Basics of Gas Field Processing Book "Full text"
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Basics of Gas Field Processing Book "Full text"
Chapter 1 - Part 1
Chapter 1 11
Basics of Oil and Gas Treatment 11
1.1 Introduction 11
1.2 Hydrocarbon preparation 11
1.3 Produced Hydrocarbon Fluids 12
1.3.1: Hydrocarbon gases 14
1.3.2: Molecular weight and apparent molecular weight 16
1.3.3: Apparent molecular weight of gas mixture 17
1.3.4: Gas Specific Gravity and Density 18
1.3.5: General Gas Law 19
1.4 Natural Gas Field Processing 19
1.5 Natural Gas Composition 21
1.6 The heating value of gases 23
1.7 Natural Gas Sampling 24
1.7.1 General Overview and introduction 24
1.7.2 Sample procedures and precautions 26
1.8 Product specifications 32
1.8.1 Natural gas 32
1.8.2 Natural-Gas Liquids 35
1.9 Physical properties of Hydrocarbon Gases 35
1.9.1 Compressibility Factor (z) 35
1.9.2 Gas density at any condition of Pressure and temperature 38
1.9.3 Gas volume at any condition of Pressure and temperature 39
1.9.4 Velocity of gas, (ft/s) 41
1.9.5 Average pipeline pressure 42
1.9.6 Viscosity of gases 43
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Chapter 1
Basics of Oil and Gas Treatment
1.1 Introduction
Oil and gas wells produce a mixture of hydrocarbon gas, condensate or oil, salt water, other gases, including nitrogen, carbon dioxide (CO2), and possibly hydrogen sulfide (H2S), and solids, including sand from the reservoir, dirt, scale, and corrosion products from the tubing.
These mixtures are very difficult to handle, meter, or transport. In addition to the difficulty, it is also unsafe and uneconomical to ship or to transport these mixtures to refineries and gas plants for processing. Further, hydrocarbon shipping tankers, oil refineries, and gas plants require certain specifications for the fluids that each receive. Also, environmental constraints exist for the safe and acceptable handling of hydrocarbon fluids and disposal of produced salt water. It is therefore necessary to process the produced fluids in the field to yield products that meet the specifications set by the customer and are safe to handle.
1.2 Hydrocarbon preparation
The goal is to produce oil that meets the purchaser’s specifications that define the maximum allowable amounts of water, salt, and sulfur. In addition to the maximum allowable value of Reid vapor pressure and maximum allowable pour point temperature.
Similarly, the gas must be processed to meet purchaser’s water vapor maximum allowable content (Water dew point), hydrocarbon dew point specifications to limit condensation during transportation, in addition to the maximum allowable content of CO2, H2S, O2, Total Sulfur, Mercaptan, Mercury, and maximum gross heating value.
The produced water must meet the regulatory requirements for disposal in the ocean if the wells are offshore, or to meet reservoir requirements for injection into an underground reservoir to avoid plugging the reservoir.
The specifications for the above requirements may include maximum oil in water content, total suspended solids to avoid formation plugging, bacteria counts, toxicity in case of offshore disposal, and oxygen content. Before discussing the industry or the technology of oil and gas processing it is best to define the characteristic, physical properties and main chemical composition of oil and gas produced.
Oil field gas processing which is the subject of this book; generally consists of two distinct categories of operations:
Separation of the gas-oil-brine well-stream into its individual phases,
Removal of impurities from the separated phases to meet sales/transportation/reinjection specifications and/or environmental regulations.
Figures 1-1 and 1-2, illustrates gas-oil separation plant, and gas flow diagram.
Fig.1-1 .Gas Oil Separation plant function and gas flow diagram.
1.3 Produced Hydrocarbon Fluids
The desirable constituents of crude oil and natural gas are hydrocarbons. These compounds range from methane (CH4) at a lower-molecular-weight end all the way up to paraffin hydrocarbons with 33 carbon atoms and poly-nuclear aromatic hydrocarbons with 20 or more carbon atoms. Natural gas is principally methane. Crude oil is principally liquid hydrocarbon having 4 and more carbon atoms.
There are a tendency to regard crude oil as a liquid and natural gas as a gas and consider production of the two phases as separate operations. However, in the reservoir, crude oil almost always contains dissolved methane and other light hydrocarbons that are released as gas when the pressure on the oil is reduced. As the gas evolves, the remaining crude-oil liquid volume decreases; this phenomenon is known as shrinkage. The gas so produced is called associated or separator gas. Shrinkage is expressed in terms of barrels of stock-tank oil per barrel of reservoir fluid. Crude oil shrinkage is the reciprocal of oil formation volume factor (FVF).
Similarly, natural gas produced from a gas reservoir may contain small amounts of heavier hydrocarbons that are separated as a liquid called condensate. Natural gas containing condensate is said to be wet. Conversely, if no condensate forms when the gas is produced to the surface, the gas is called dry.
A hydrocarbon constituent range or a spectrum of well fluids usually produced are summarized in table 1-1 (as noted by McCain (1973)). The type of fluid produced depends on the phase diagram of the reservoir fluid and the reservoir temperature and pressure, as will be discussed later in phase behavior of natural gas.
Notes.
Separator may be a slug catcher, free water knock out drum, two phase separator, or gun barrel.
Oil and water are separated and undergoes further treatment processes not in the scope of this book.
Fig.1-2. An example of gas flow diagram
Figure 1-3 depicts a typical gas-oil separation sequence (including incidental water removal). Table 1-1 lists the five common types of wellstream fluids and summarizes typical yields and liquid properties.
Table. 1-1. Petroleum fluid spectrum (after McCain, 1990)
When crude oil is separated from its associated gas during production, the total gas evolved while reducing the oil to atmospheric pressure divided by the volume of the remaining crude oil is called gas-oil ratio or GOR. The GOR is expressed as the total standard cubic feet of gas evolved per 60 0F barrel of stock-tank or atmospheric-pressure oil (scf/bsto). SI units are standard cubic meters of gas per cubic meter of 15 0C oil (metric units).
The total GOR depends on the number of stages used in the separation sequence, as well as the operating pressure of each stage. For three or more stages, the GOR approaches a limiting value. Optimization of the separation sequence usually involves either maximizing crude-oil yield or minimizing recompression horsepower as well be explained briefly in chapter 3.
For wet natural gas, the liquid content is given in barrels of condensate per million standard cubic feet of gas (bbl/MMscf) or in U.S. gallons of condensate per thousand standard cubic feet (GPM).
The various types of produced hydrocarbons have been described using GOR (McCain, 1973- and after McCain, 1990. Table 1-1) or composition (Gould and McDonald, 1979 – Table 1-2).
Of course the gas yield depends on the relative amounts of the various hydrocarbons present, and perhaps Penick (1983) summarizes these relationships best in figure 1-3 ( originally drawn by W.H. Speaker, Jr.).
Allen (1952) emphasizes that classifications like Table 1-1. are an oversimplification in the sense that GOR does not always reflect the condition of the fluid in the reservoir. For example, in the GOR range of 3000 – 7000 scf/bbl the reservoir fluid may be a liquid or a denese fluid (highly-compressed gas) but still be a black oil on the surface. If the fluid is a denese fluid at reservoir temperature and pressure, then any liquid produced is usually defined as condensate. If the reservoir fluid is a liquid, then the produced liquid is called volatile oil. The difference in these two cases is explained in chapter 2.
Table -1-2 . Well-Fluid Type (Gould and McDonald, 1979)
1.3.1: Hydrocarbon gases
Most of compounds in crude oil and natural gas consist of molecules made up of hydrogen and carbon, therefore these types of compounds are called hydrocarbon.
The smallest hydrocarbon molecule is Methane (CH4) which consists of one atom of Carbon and four atoms of hydrogen. It may be abbreviated as C1 since it consisted from only one carbon atom. Next compound is Ethane (C2H6) abbreviated as C2, and so on Propane (C3H8), Butane (C4H10)...etc.
Hydrocarbon gases are C1:C4), with the increase of carbon number, liquid volatile hydrocarbon is found (e.g. Pentane C5 is the first liquid hydrocarbon at standard conditions).
Fig. 1-3. Typical Reservoir composition (Penick, 1983).
Fig.1-4. Stage separation (Gas-oil separator train).
1.3.2: Molecular weight and apparent molecular weight
The molecular weight of a compound is the sum of the atomic weight of the various atoms making up that compound. The Mole is the unit of measurements for the amount of substance, the number of moles is defined as follows:
Mole = Weight/(Molecular weight) (Eq. 1-1)
Expressed as n = m/M (Eq. 1-2)
or, in units as, lb-mole = lb/(lb/lb-mole) (Eq. 1-3)
Table 1-3 Physical constants of light hydrocarbons and some inorganic gases. Adapted from GPSA, Engineering Data Book, chapter 23.
Example 1.1: Methane molecule consists of one carbon atom with atomic weight =12 and 4 hydrogen atoms with atomic weight = 1 each. Molecular weight for Methane (CH4) = (1 × 12) + (4 × 1) = 16 lb/lb-mole. Similarly, Ethane (C2H6) molecular weight = (2 × 12) + (6 × 1) = 30 lb/lb-mole.
Hydrocarbon up to four carbon atoms are gases at room temperature and atmospheric pressure. Reducing the gas temperature and/or increasing the pressure will condense the hydrocarbon gas to a liquid phase. By the increase of carbon atoms in hydrocarbon molecules, consequently the increase in molecular weight, the boiling point increases and a solid hydrocarbon is found at high molecular weight.
Physical constants of light hydrocarbon and some inorganic gases are listed in Table 1-3.
1.3.3: Apparent molecular weight of gas mixture
For compounds, the term molecular weight is used, while, for hydrocarbon mixture the term apparent molecular weight is commonly used. Apparent molecular weight is defined as the sum of the products of the mole fractions of each component times the molecular weight of that component. As shown in Eq. 1-4
MW= ∑▒ Yi (MW)i (Eq. 1-4)
where
yi =molecular fraction of ith component,
MWi =molecular weight of ith component,
Ʃyi =1.
Example 1.2: Determine the apparent molecular weight for the gas mixture in Table 1-4:
Table 1-4 Gas mixture for Example 1-2
Solution: Using Table 1-3 & Equation 1-4
MW= ∑▒ Yi (MW)i
MW = (Mole Fraction of component 1 × MW of component 1) + (Mole Fraction of component 2 × MW of component 2) + (Mole Fraction of component 3 × MW of component 3) +…etc.
The following table can be made:
Table 1-5 Solution of Example 1-2
The apparent molecular weight is 21.36
1.3.4: Gas Specific Gravity and Density
The density of a gas is defined as the mass per unit volume as follows
Density = mass / volume (Eq. 1-5)
The specific gravity of a gas is the ratio of the density of the gas to the density of air at standard conditions of temperature and pressure.
S = (ρ(gas))/(ρ(air)) (Eq. 1-6)
Where
ρ(gas) ρg = density of gas
ρ(air) ρair = density of air
Both densities must be computed at the same pressure and temperature, usually at standard conditions.
It may be related to the molecular weight by Equation 1-7
S = (MW(gas))/(MW(air)) (Eq. 1-7)
Since molecular weight of air is 28.96 (table 1-3)
Specific gravity of gas S = (MW(gas))/28.96 (Eq. 1-8)
Example 1-3: Determine the specific gravity of the gas mixture in example 1-2.
Solution:
Apparent molecular weight of gas mixture is 21.36
Gas specific gravity = 21.36/28.96 = 0.7376
Since the gas is a compressible fluid, its density varies with temperature and pressure, calculating the gas density at a certain pressure and temperature will be explained after discussing the general gas law and gas compressibility factor.
1.3.5: General Gas Law
The general (Ideal) Gas equation, or the Perfect Gas Equation, is stated as follows:
PV = nRT (Eq. 1-9)
Where
P = gas pressure, psia
V = gas volume, ft3
n = number of lb moles of gas (mass/molecular weight)
R = universal gas constant, psia ft3/lb mole OR
T = gas temperature, OR (OR = 460 + OF)
The universal gas constant R is equal to 10.73 psia ft3/lb mole OR in field units.
Equation (1-9) is valid up to pressures of about 60 psia. As pressure increases above this level, its accuracy becomes less and the system should be considered a non-ideal gas equation of state.
PV = znRT (Eq. 1-10)
Where
z = gas compressibility factor.
1.4 Natural Gas Field Processing
The main constituent of natural gas is methane, desirable as a primary fuel. Sales gas also contains smaller amounts of the heavier hydrocarbons listed in table 1- 6. Often a portion of the heavier hydrocarbon can be recovered profitably in a field-gas processing plant as one or more liquid products. These liquefilable components ( or condensate) may be recovered as a single liquid stream that is transported to a separate plant for fractionation into stable products. Alternatively, in very large field units, fractionation is performed in the field. Common natural gas liquid (NGL) products are summarized below.
Table. 1- 6. Common natural gas liquid (NGL) products nad uses.
The permanet gases occuring in natural gas include nitrogen, helium, argon, hydrogen, and oxygen. Most natural gases contains some nitrogen, and a few have 30 mole percent or more. Nitrogen lowers the heat of combustion of the gas. Since natural gas is usually sold on the basis of energy content with a fixed minimum heating value, the nitrogen content is limited to fairly low amounts in commercially salable gas. The removal of nitrogen require expensive cryogenic processing, so too high a nitrogen content may render a gas unusable.
Many natural gases contain a few hundredth of a percent of helium. Helium has no deleterious effect other than lack of heating value. Separated helium is very valuable and, in the U.S. for example, suffficient lawsuits have awarded royalty payments for helium that could have been recovered economically from sales gases. Analysis also reported occasional small amounts of oxygen, as well as argon and hydrogen. Chromatographic analyses may lump all the inert gases as nitrogen.
Hydrogen sulfide and carbon dioxide are found in many Natural gases and may occur in very high percentages. In fact, essentially pure carbon dioxide is produced, dehydrated, and pipelined for CO2 floods in the enhanced recovery of crude oil. Hydrogen sulfide and carbon dioxide are referred to as acid gases because they dissociate upon solution in water to form acidic solutions. Hydrogen sulfide is very toxic and corrosive, while carbon dioxide is corrosive.
A natural gas containing no hydrogen sulfide is said to be sweet. Conversely, a sour gas containing hydrogen sulfide. Strictly speaking, “sweet” and “sour” refer to both acid gases (CO2 and H2S) but are usually applied to hydrogen sulfide alone.
Sulfur compounds, other than H2S, are present in minute amounts and can affect field processing. These compounds tend to concentrate in the condensate and sometimes require treating (or sweetening) of the liquid products.
Removal of hydrogen sulfide to very low content (4 ppmv or 1/4 grain/100 scf) is required in the field. Carbon dioxide can be tolerated to much higher levels, say 1-2%, as long as the heating value of the sales gas is satisfactory.
There are many so-called “treating” processes for sweetening natural gas. These processes are either batch, reactant-discarded processes for removing low amounts of hydrogen sulfide or continuous solvent-regenerated processes for large amounts. Batch processes are used when the consumable-chemical cost is not prohibitive. The principal continuous, solvent regenerated treating processes use water solution of chemical solvents, typically alkanolamines. Other sweetening such as physical solvents, mixed physical-chemical solvents, or membranes may be more economical in some cases.
Because these processes also remove appreciable amounts of other sulfur compounds and/or carbon dioxide, many gas streams need no further processing to meet total sulfur and acid gas specifications.
Hydrogen sulfide is an extremely toxic substance. Fortunately, the familiar sulfurous smell can be detected at concentration less than 1 ppmv. However, extended exposure at higher levels deadens the sense of smell, so that odor alone cannot be a reliable detector.
Water or brine is undoubtedly present in many gas reservoirs but usually is not entrained to the surface. If free liquid water or brine is produced, a wellhead knockout drum vessel, or separator, is needed to prevent the water entering the gathering lines. At high pressure and low temperature, natural gas and free liquid water form solid hydrates capable of plugging flow lines. Produced water also may contain methanol and/or corrosion inhibitors that have been injected in the well string.
Produced gas is regarded as being saturated with water vapor at the wellhead conditions of temperature and pressure, even if no liquid water is produced. Associated gas is regarded as being saturated with water vapor at the outlet from the gas-oil separator in which it is produced. Water condensed downstream of the wellhead or gas-oil separator will be essentially pure (fresh) water rather than saline brine.
Water vapor is a contaminant that must be removed by proper processing of the natural gas stream. Gas transmission lines often specify water content of 7 lb/MMscf (or a water dew point of 32 0F or less). Triethylene glycol (TEG) is almost always used for such applications. Cryogenic plants require “bone-dry” gas (water dew point as low as -150 0F). Solid desiccants dehydration is typically used for dew points below – 25 0F.
Formation solids are not produced with most natural gas. Nevertheless, solids are sometimes separated from the process plant inlet separator. Usually these solids are mainly mill scale and rust from pipe wall, along with iron sulfide (for sour gas). The solids interfere with treating and dehydration processes and should be removed in a scrubber, filter, or possibly a filter-coalescer separator.
Very deep, high temperature, high pressure sour gas, may contain solid sulfur. Sulfur is inert but must be properly separated to prevent downstream processing problems.
Mercury has been detected in natural gas streams, in concentrations from approximately 1 ppbw to 230 ppmw. Mercury usually causes no processing problems but has caused corrosion of aluminum tubes in heat exchangers. Removal using sulfur-impregnated activated carbon, sulfur beds, or molecular sieves has been suggested.
Recently, arsenic compounds were reported present in natural gas produced from some fields. The discovery was made indirectly by the detection of white powder on regulators in gas company. The arsenic will be removed using a vertical copper-zinc adsorbent bed.
Methanol (to prevent hydrate formation) and corrosion inhibitors are sometimes added downhole and so may be present in natural gas at the wellhead. Any compressed gas will contain some entrained lubricating oils. Field processing may also introduce additional contaminants such as glycols and amines. Mixtures of these liquids with the previously listed solid contaminants are called sludge.
1.5 Natural Gas Composition
The gas analyses shown in table 1-7, span the composition ranges normally encountered. These analyses are typical of the data furnished to the designer of surface processing equipment. As previously mentioned, the heavier hydrocarbon in these gases are regarded as a recoverable liquids. The amount of potentially recoverable liquid is expressed as gallons liquid at 60 0F, if totally condensed, per 1000 standard cubic feet of gas (so called GPM, not to be confused with gallons per minute). A gas is termed lean or rich as follows:
Lean < 2.5 GPM
Moderately-Rich 2.5-5 GPM
Very Rich > 5 GPM
The above classification is based on ethane and heavier hydrocarbons (C2+) because ethane is sometimes regarded as a desirable feed for petrochemical process and can be recovered as a liquid in expander-type gas processing plants. If ethane is not considered as a valuable liquid component, the GPM can be based on propane and heavier hydrocarbons (C3+).
Table. 1-7. Typical gas analyses (Mol Percent).
Example 1-4. Confirm the GPM given in table 1-7. For the first natural gas stream.
Solution: First find the scf/gal of the C7+ components
scf C7+/ gal liq C7+ = (SG C7+)*(lb H2O/gal)*(1/Mol.Wt.)*(scf/lb mol) (Eq. 1-11)
= (0.803) * (8.334) * (1/172) (379.5) = 14.77
(379.5 = n and is a constant of (scf/lb mol) derived from PV=nRT at standard condition)
The table that follows details the calculations (which are most conveniently done with a spread- sheet program)
Table. 1-8. Solution of example 1-4.
(a) From Table 1-3.
(b) Column 3/ column 4
GPM is approximately = 0.3 * [100 – (C1 + N2+CO2)].
The weak link in gas analysis is often the composition of the C6+ or C7+ portion of the gas. For many purposes the small amount of the C6+ material renders its characterization unimportant. One important exception is when a full wellstream gas is to be transported in a pipeline over a long distance, such as from an offshore platform. Condensation of liquids in the line will cause a large pressure drop that must be anticipated if adequate platform compression is to be furnished and the proper pipe diameter selected. Very accurate characterization of the C6+ is needed for hydrocarbon (HC) dew points prediction, otherwise the HC can be determined in laboratory.
1.6 The heating value of gases
The heating value of a gas is expressed in Btu/ft3. It represents the quantity of heat in Btu (British Thermal Unit) generated by the complete combustion of one cubic foot of the gas with air at constant pressure (1 atmosphere = 14.7 psia) and at a fixed temperature of 60 0F.
Hydrogen in the fuel burns to water and when the flue gases are cooled to 60°F, the physical state — either vapor or liquid — of this water must be assumed. So the latent heat of vaporization of the water may or may not be considered to be part of the heating value. The result is two definitions for the heating value. The higher or gross heating value, HHV, includes the heat of condensation and the lower or net heating value, LHV, assumes the water remains in the vapor state.
For gas mixture the heating value is calculated as follows:
H = Ʃ xi Hi Eq. 1-12
Example 1-5: Calculate the heating value of gas mixture of Example 1-2
Table 1-9 Solution of Example 1-5
From table 1-9 the Gross calorific value HHV = 1246 Btu/ft3
The higher, ideal, dry heating value of sweet natural gas at 60°F and 760 mm Hg may be calculated with the following equation:
HHV=1568.72 × SG – 2524.88 × XCO2 – 1658.37 × XN2 +141.05 (Eq. 1- 13)
Applying Eq 1-13 for Example 1-5
The apparent molecular weight= 21.36
Gas Specific gravity = 21.36/28.96 = 0.738
HHV = 1568.72 × 0.738 – 2524.88 × 0.015 – 1658.37 × 0.01 +141.05 = 1244 Btu/ft3
1.7 Natural Gas Sampling
1.7.1 General Overview and introduction
The purpose of natural gas sampling is to secure a representative sample of the flowing gas stream for a specific period of time. Naturally, the more often the sampling system samples the flowing stream, the more likely it is to be truly representative of a stream with varying composition. Refer to figure 1-5.
Sampling systems consist of numerous components and must include some key elements including a sample probe, any necessary connecting tubing, sample containers or sample valves and appropriate heat-tracing and insulation. For online gas chromatographs, or on-stream analyzers, the sampling system ends at the injection valve on the inlet of the g.c.
The sample may be collected on a spot, composite or continuous basis.
Spot sampling simply means that a technician manually collects a sample directly from the stream at scheduled intervals or as needed.
Composite samples are usually collected on a weekly or monthly basis. Composite sampling systems should grab small samples on a flow proportional basis, then inject them into the composite sampling cylinder. There are composite sampling systems that work on a simple time cycle (time proportional sampling), but they are not recommended, especially if they continue to sample even when flow has stopped. If a time proportional system is already in service, it must be equipped with a flow switch or similar device to ensure that when flow stops, sampling will stop.
Continuous systems provide a steady flow of sample through a sample loop that passes near a composite sampler or on stream analyzer. In the case of an on-line chromatograph, the injection valve of the g.c. is able to admit and distribute a sample from the flowing loop that is representative of the flow in the main line. Sample rate flow loops must be carefully sized and generally should operate at velocities around 5 ft/sec., but this may vary if the sample loop is exceptionally long (over 100 feet).
The sample point is usually located downstream of the meter run and must be remote from severe flow disturbances such as control valves and orifice plates by at least five nominal pipe diameters. For flowing streams that are not near their hydrocarbon dew point, the probe should be positioned either upstream or downstream of the meter tube, and at least 5 pipe diameters downstream of any flow disturbing elements, such as elbows, swirl generators, headers, valves and tees. If the sample source is at or near its hydrocarbon dewpoint, some research has indicated that the probe should be located at least 8 pipe diameters downstream of any flow disturbance, including an orifice meter. The sample point must not be installed within the upstream or downstream engineered sections of the meter tube, since the fitting and probe could produce disturbances in the flow profile going through the meter in the run. The probe installed in the sample point extends into the center 1/3 of the internal diameter of the meter run to ensure no heavy materials or contaminants migrating along the pipe walls are allowed to contaminate the sample. Note that for large diameter pipelines, the probe never needs to be longer than 10 inches. The probe is equipped with an outlet valve to allow the system to be shut in when no sampling is being performed or to perform maintenance on downstream equipment in composite or continuous sampling systems.
The tubing connecting the sample probe to the downstream sample system(s) should be internally clean, as short as practical (usually 6 to 24 inches maximum) and made of either nylon or stainless steel. Stainless steel is actually preferred due to its strength and flexibility and resistance to melting and/or sharp edges, but nylon is not porous and when used safely, can also give good analytical results. Teflon, carbon steel, plastic tubing and many other materials do not perform well.
Care must be taken to insure there are no leaks in the sampling system. Typically, if a leak occurs, smaller molecules tend to escape preferentially and create a bias in analytical results. If the leak is large, there may be enough cooling to produce condensation in the sample system and cause the samples to be very non-representative.
Note that whenever the sample line is operating in ambient temperatures below the flowing temperature of the stream, the line may need to be heat-traced and insulated. If the ambient temperature is lower than the dewpoint temperature of the flowing stream, heat tracing and insulation are required.
Be sure that the heat tracing is properly and safely done, using electrically limited tracing meeting appropriate electrical codes for the area classification (typically Class I Group D Division I or II).
Realize that the dewpoint of a gas stream is absolutely critical to accurate sampling. If any component in the sampling system causes the temperature of even a portion of the gas stream being sampled to cool to or below the hydrocarbon dewpoint, the sample will be depleted of heavy components and can no longer be truly representative of the stream. Note that the Btu content in this situation is not always too low. If the sample system continues to condense heavy components for an extended amount of time, accumulations may reach the point that liquid droplets enter the sample and actually cause the indicted Btu content and calculated relative density (specific gravity) to be too high.
The sample cylinders used in spot sampling should be stainless steel, single cavity cylinders. Single cavity cylinders are recommended, due to the difficulty of fully cleaning piston cylinders between uses. Residue that may remain in the piston cylinders and their seals may produce incorrect analyses. The cylinders should be equipped with standard design sample valves that are screw open or closed (not 1/4 turn ball valves) and have a flow passage of approximately 1/8 inch diameter.
It may be convenient to use piston (constant pressure) cylinders in composite sampling systems, since you can easily see that the system is working or not working as the level indicator moves. If constant pressure/piston cylinders are used and oil or grease contamination is present in the system, they must be disassembled between uses, carefully cleaned and then the seal rings must be replaced if the cylinder is expected to provide representative samples.
Note that if sampling is being performed to determine the levels of volatile or reactive contaminants, such as H2S, the cylinder may need to be lined with an epoxy/phenolic lining.
Even then, particularly reactive materials, such as H2S or ethyl mercaptan are likely to be lost prior to analysis unless the sample is collected on-site and analyzed immediately.
Even a few minutes delay can reduce detectable levels of reactive materials. Shipping a sample to a remote lab and delaying analysis beyond a couple of hours will essentially ensure that the indicated levels of the reactive/volatile material will be too low or perhaps not detectable at all.
The two spot sampling methods that are most recommended are the fill and empty method and the helium pop method. The displacement methods also performed reasonably well during the recent API research studies.
The fill and empty method requires that the cylinder be equipped with a “pigtail” following the sample cylinder outlet valve. While leaving the sample cylinder inlet and outlet valves open, the probe outlet valve and the valve at the end of the pigtail are cycled to alternately fill and empty the sample cylinder. The pigtail ensures that the heat of compression created when the sample cylinder is filled more than offsets the Joule-Thomson cooling produced when the sample cylinder is depressurized. It does this by insuring the maximum pressure drop while depressurizing the sample cylinder is far removed from cylinder itself, at the end of the pigtail. The ability of the fill and empty procedure to actually elevate the temperature of the sampling cylinder above the flowing temperature of the stream being sampled during many operating conditions makes this method the most desirable when the ambient temperature is at or near the hydrocarbon dewpoint of the stream.
The pigtail should be approximately 1/4 inch tubing and be at least 36 inches in length, although it may be coiled to make the apparatus easier to handle. The coils should not touch one another, otherwise the heat loss at the end of the pigtail may be transferred quickly across the coils to the sample cylinder.
There should be another sample valve, similar to the sample cylinder valves on the outlet of the pigtail. The flow passage through this valve must not be larger than the passage through the cylinder valves. Refer to API Chapter 14.1 for the detailed procedures for performing the fill and empty method and to either API Chapter 14.1 or GPA 2166 for the number of fill and empty purge cycles required at various line pressures.
1.7.2 Sample procedures and precautions
Collection of a truly representative samples is not a simple job as it looks like. A moment’s thought will confirm that any subsequent design, operating, or investment decision can be no better than the prior sampling and analyses.
In spite of its obvious importance, natural-gas sampling is seldom done well. Natural gas, flowing with accompanying condensate, can not be properly sampled by withdrawal of a portion through a sample tap. Reflection of the complex natural of two phase flow (with its varying flow patterns and liquid-holdup phenomenon, along with velocity gradients in the fluid) should convince one of the impossibility.
In the ideal case, a separator should be set up at the wellhead to collect condensate and meter each phase. Subsequent analysis of the separated gas and liquid allows recombination to obtain the wellstream analysis. This is the reason well-test separators are even installed offshore where space and weight are extremely expensive.
Poor sampling is generally caused by:
Ignorance of its economic and technical importance.
Inadequate training of personnel involved.
Ignorance and/or unwillingness to follow standard procedures.
While sampling in the field do not:
Use dirty cylinders that contain previous samples.
Sample when pressure and temperature are not stable
Sample when separator pressure is different from sample pressure.
Dilute sample with air.
Fail to identify sample completely.
Sampling hydrocarbon fluids can be hazardous. Every body involved in sampling must be familiar with and follow safe practices while handling flammable fluids under pressure. All sample containers must meet the specification to handle sampled fluids, and must be clearly labeled.
Do not fill liquid sample containers completely full. A 150 ml, stainless-steel container was filled full with gasoline and closed at 0 psig and 61 0F. When heated to 105 0F, the sample pressure was 3100 psig!.
1.7.2.1 Gas sampling
GPA standard 2166-86 indicates that good samples can be obtained by all 8 approved methods provided extreme care is taken, however the sampling project report ranked the accuracy of the eight methods in the following order of decreasing accuracy (refer to fig. 1-5):
Water displacement
Continuous purge
Purge and fill
Glycol displacement
Reduced pressure
Floating piston cylinder
Helium filled to 5 psig
Evacuated container
In practice purge and fill is the most popular method while the floating piston cylinder is finding increased use because composite samples cam be obtained.
Several problems can arise during sampling of natural gas:
Condensation of hydrocarbons due to temperature and pressure changes during sampling.
Entrainment of liquid droplets and mist
Sample constituents can react with sample container.
Some sample components may dissolve in the displacement media.
Subsequent handling of the sample is therefore, very important to assure the relatively small amounts of condensable components remain in the gas phase and are removed to the chromatograph or mass spectrometer.
In a natural gas, hydrogen sulfide in small amounts may be overlooked entirely, creating a future problem in design and operation. Hydrogen sulfide reacts rapidly with carbon steel and may disappear from the sample altogether. A stainless-steel sample cylinder should be used if the presence of H2S is suspected. Even austenitic stainless-steel will absorb a small amount of H2S. GPA Standard 2261 recommends that the gas to be analyzed at its source for hydrogen sulfide content less than 3% by GPA Standard method 2377. Teflon-lined cylinders have also been used successfully.
The following precautions are recommended:
Use a pre-filter and/or “knockout” trap on the sample system just downstream from the source valve. This is mandatory on “wet” gas samples.
Use heated sample lines to prevent condensation from low ambient temperature during sampling.
Keep sample lines as short as possible. This should be done for all sample grabbing, wet or dry.
Clean, dry, and evacuate sample cylinders before taking to the field. This prevents possible liquid carryover from the previous sample.
Purge sample cylinder carefully.
Empty cylinders through a non-metallic pigtail on the outlet for the expansion of pressured gas to the atmosphere.
For samples requiring calorimetric heating value determination and chromatographic analysis, use stainless steel or carbon steel, DOT3A and DOT3AA, cylinders, 300 cu.in, volume. Smaller cylinders may be used for GC analysis only; however, a calorimeter Btu determination provides good confirmation of the chromatographic analysis. Calculated Btu from the analysis should not be over +/- 3 Btu from the measured value.
Keep sample cylinders upright while filling. With the purge line at the bottom. Do not lay the cylinders on the ground.
Use line probe that extends at least a third of the way into the line of sampling. This prevents the entry condensate and other contaminants into the sampling line.
Sampling by liquid displacement has been used in the past and still followed by some organizations. It is not recommended for gases or liquids containing acid-gases such as CO2 and H2S. These components dissolve in the liquid, usually water, and are lost. Therefore the resulting analysis is not representative.
Continuous sampling
The following procedures and precautions are recommended for continuous or composite sampling of natural gas:
Sample point should ”see” center one-third of the pipeline in an area of good velocity with minimum turbulence.
Sample probe, equipped with full open ball or gate valve, must be kept away from pipeline fittings and orifice plates. The probe may be bevel or flat cut at end, but must be kept clear of free liquids and aerosols. Bevel may face upstream or downstream.
Probe construction: stationary or permanent, manually insertable or automatic insertion type probe s may be used. The probe should be stainless steel so that it will not react with the sampled gas.
Sampler hook-up and manifold: The sampler should be mounted above the sample point as shown in figure 1- 6. Other precautions are:
a. Line from probe to sampler should slope down to let any free liquid to drain back into pipeline.
b. Never sample a dead-end line.
c. Any leaks in the line from the sampler to the cylinder will “lose” light ends preferentially.
d. Any filters, drips, or regulators between the probe and sampler will invalidate the sample.
Sampler: Should take a composite sample in the same way as an operator would take a spot sample. If the pipeline flow rate varies, the sampler should be actuated proportional to the flow. The sampler should be able to pump the sample into the cylinder and purge itself before pumping each new “bite”.
Sample containers
Proper cleaning and inspection of sample cylinders cannot be overemphasized. Several methods are available:
Volatile solvent and air dry
Steam clean and air dry
Evacuation
Evacuation and fill with 5 psig helium
Correct procedures must be followed rigorously in every detail.
FIG. 1-5. Fill and Empty Sampling Method, and helium pop sampling method.
Fig.1-6 .Gas automatic or continuous sampler.
1.7.2.2 Natural Gas Liquid Sampling
Utilizing Floating Piston Cylinders
Liquid sampling requires special precautions to accumulate and transfer representative samples. Pressure in the sample cylinder and/or accumulator must be maintained at 1.5 times the product vapor pressure. Maximum product vapor pressure should be determined using the highest ambient temperature or flowing temperature (whichever is highest) to determine the minimum pre-charge pressure. A method to break up stratification must be provided prior to transfer of the sample to another container and laboratory analysis. Maintaining the appropriate pressure and mixing the sample can be satisfied by using floating piston sample cylinders with mixers (Fig. 1-7). A “rattle ball” or agitator may be used in place of the mixing rod shown. The floating piston cylinders are precharged on one end with an inert gas at a pressure 1.5 times above product vapor pressure. This prevents sample vaporization, which could result in erroneous analysis. This design also provides a compressible inert gas cushion to allow for thermal expansion of the liquid. A pressure relief valve is needed, but should it discharge, the integrity of the sample will be lost.
Liquid sample cylinders shall not be filled over 80 percent full.
FIG. 1-7. Liquid Sample Container
Samples are acquired through a sample probe inserted into the center third of the flowing stream. The probe should be mounted in the top or side of the line. Continuous samplers should use a continuous flowing sample loop (speed loop) or a probe mounted sample pump to ensure the most current sample is always added to the sample container. Speed loops must have a driving device such as an orifice, differential pump, available pressure drop, or “scoop” probes. The driving device should be sized to provide a complete exchange of liquid in the sample loop once per minute. The sample pump must be set to gather flow proportional samples to ensure a true representative sample is obtained. If flowing pressures are higher than the sample accumulator pre-charge pressure, then the sample pump must prevent “free-flowing” of product into the sample container. See Fig. 1-8 for an example continuous sampling application. Speed loop lines may require insulation when cold ambient temperatures have a significant effect on viscosity. The product in the sample container must be thoroughly mixed before being transferred to a transport cylinder.
Details and alternative methods for obtaining liquid samples are found in GPA 2174.
Fig. 1-8. Continuous Sampler (Automatic sampler)
Procedures and precautions for liquid sampling
A GPA Work group conducted a cooperative natural gas liquid sampling project for future revision of GPA Standard 2174. The following four sampling methods were judged acceptable:
Floating Piston Cylinder.
Water displacement (total removal – 80% hydrocarbons/20%displaced outage)
Water displacement (partial removal – 70% hydrocarbons/20%displaced outage/10%water remaining in cylinder)
Ethylene glycol displacement (total removal – 80% hydrocarbons/20%displaced outage)
The following precautions are recommended for liquid sampling while using the floating piston cylinder:
Pressure on the backside piston should be higher than the line pressure at the beginning of operation
Slowly bleed pressure down to enter sample, and maintain pressures on each side of the piston at nearly equal levels with just enough difference to move the pistons
Do not bleed the pressure off the backside of the piston after sample has been taken, this will flash some of the liquid into gas. Then you will need to resample
Do not fill cylinder 100% full with sample. Leave at least 25% as a pressure buffer to take care of ambient temperature fluctuations.
The floating piston cylinder can be used for “wet” gas sampling with excellent results in getting representative samples. Procedures are as follows:
Pressure up back of cylinder with line gas to full line pressure.
Connect cylinder to source (use of a line probe is mandatory) and open valves to cylinder.
Fill cylinders as you would in taking liquid sample. Note: It is extremely important to avoid any pressure drops. The pressure differences should not exceed 2 psi maximum. This will prevent flashing of heavy components which may be in aerosol form. Sample should be slowly entered into cylinder.
1.8 Product specifications
The objective of field processing is to provide transportable and/or salable fluids. There are two main products: natural gas and condensate (raw mix) or natural gas liquid (NGL).
1.8.1 Natural gas
Table 1-10 lists typical natural gas pipeline specifications listed in a typical sales gas contract. Theses specifications are fixed by negotiation between seller and buyer and vary from case to case. Not all sales gases will have all the specifications shown in the following items.
Table. 1- 10. Natural gas pipeline specifications.
Table. 1-11. Is a typical sales gas/ pipeline contract between two middle east companies.
Table. 1-11. Typical sales gas/ pipeline contract.
Wobbe Index is some times used in sales gas specification. Wobbe Index = [Gross heating value / (Sp.Gr)0.5].
The most important specifications are: water content (water dewpoint), H2S content, and gross heating value. The following table summarizes the effect of water, acid gases, and liquid hydrocarbons in sales gases:
Table.1-13. Water dew point and water content relation. (ANSI/CGA G-7.1 -1989 document).
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Note that both water and hydrogen sulfide must be removed to very low concentrations. Heating value is more complex (specification usually from 950 to 1200 Btu/scf).
Refer to table 1-3. For heating value of hydrocarbon gases. (Complete physical properties of hydrocarbon gases are listed in GPSA, Engineering data book, chapter 23).
From table 1-3, we can realize that; the most abundant component, methane, has a relatively low heating value (1010 Btu/scf). Therefore, by itself methane cannot always fulfill the minimum heating value requirement when inert gases (nitrogen and/or carbon dioxide) are present. However, enough heavier hydrocarbons are usually present to provide the required heating value, even with condensate recovery.
When there is a maximum heating value or hydrocarbon dew point specifications, some of the heavier hydrocarbons constituents may have to be removed as condensate. Refrigeration to (-30 0F reduces the heating value) how much, of course, depends on the gas composition and pressure. Further reduction requires cryogenic processing. The remaining specifications are met by suitable processing. These processes are the subjects of the later chapters.
Fig. 1-9 Moisture Content Nomograph for Gases
1.8.2 Natural-Gas Liquids
Hydrocarbon condensate recovered from natural gas may be shipped without further processing or stabilized to produce a safely-transportable liquid. In the case of raw condensate, there are no particular specifications for the product other than the process requirements. Stabilized liquid, on the other hand, will generally have a vapor pressure specifications, since the product will be injected into a pipeline or transport pressure vessel which has a definite pressure limitations.
Natural-gas liquid products are prepared by fractionation of the raw make into appropriate products, either at the field processing site or, perhaps more commonly, at a large central facility. In any case, product specifications are not so typical as those of sales gas, but depends heavily on the particular contract. Liquid product specifications generally include composition, vapor pressure, water content, and sulfur content.
1.9 Physical properties of Hydrocarbon Gases
1.9.1 Compressibility Factor (z)
The Compressibility factor, Z is a dimensionless parameter less than 1.00 that represents the deviation of a real gas from an ideal gas. Hence it is also referred to as the gas deviation factor. At low pressures and temperatures Z is nearly equal to 1.00 whereas at higher pressures and temperatures it may range between 0.75 and 0.90. The actual value of Z at any temperature and pressure must be calculated taking into account the composition of the gas and its critical temperature and pressure. Several graphical and analytical methods are available to calculate Z. Among these, the Standing-Katz, and CNGA methods are quite popular. The critical temperature and the critical pressure of a gas are important parameters that affect the compressibility factor and are defined as follows.
The critical temperature of a pure gas is that temperature above which the gas cannot be compressed into a liquid, however much the pressure. The critical pressure is the minimum pressure required at the critical temperature of the gas to compress it into a liquid.
As an example, consider pure methane gas with a critical temperature of 343 0R and critical pressure of 666 psia (Table 1-3).
The reduced temperature of a gas is defined as the ratio of the gas temperature to its critical temperature, both being expressed in absolute units (0R). It is therefore a dimensionless number.
Similarly, the reduced pressure is a dimensionless number defined as the ratio of the absolute pressure of gas to its critical pressure.
Therefore we can state the following:
Tr = T/Tc (Eq. 1-14)
Pr = P/Pc (Eq. 1-15)
Where
P = pressure of gas, psia
T = temperature of gas, 0R
Tr = reduced temperature, dimensionless
Pr = reduced pressure, dimensionless
Tc = critical temperature, 0R
Pc = critical pressure, psia
Example1-6: Using the preceding equations, the reduced temperature and reduced pressure of a sample of methane gas at 70 0F and 1200 psia pressure can be calculated as follows
Tr = (70 +460) / 343 =1.5
Pr = 1200/666 = 1.8
For natural gas mixtures, the terms pseudo-critical temperature and pseudo-critical pressure are used. The calculation methodology will be explained shortly. Similarly we can calculate the pseudo-reduced temperature and pseudo-reduced pressure of a natural gas mixture, knowing its pseudo-critical temperature and pseudo-critical pressure.
The Standing-Katz chart, Fig. 1.10 can be used to determine the compressibility factor of a gas at any temperature and pressure, once the reduced pressure and temperature are calculated knowing the critical properties.
Pseudo-critical properties allow one to evaluate gas mixtures. Equations (1-16) and (1-17) can be used to calculate the pseudo-critical properties for gas mixtures:
P’c = Ʃ yi Pci (Eq. 1-16)
T’c = Ʃ yi Tci (Eq. 1-17)
where
P’c =pseudo-critical pressure,
T’c =pseudo-critical temperature,
Pci =critical pressure at component i, psia
Tci =critical temperature at component i, 0R
Yi =mole fraction of each component in the mixture,
Ʃ yi =1.
Example 1-7:
Calculate the Compressibility factor for the following Gas mixture at 1000F and 800 psig:
Table 1-14 for Example 1-7.
Using Equation 1-16 and 1-17
T`r = (100+460)/464.5 =1.2
P`r = (800+14.7)/659.4 = 1.23
From fig.1-10. Compressibility factor is approximately, z= 0.72
Figure 1-10 Compressibility Factor For lean sweet natural gas (Surface Production Operations).
(More graphs for compressibility factors and acid gas corrections are available in GPSA, data book)
Calculating the compressibility factor for example 1-6, of the gas at 70 0F and 1200 psia, using Standing-Katz chart, fig. 1-10. Z = 0.83 approximately. For ) Tr = 1.5 , Pr = 1.
.
Another analytical method of calculating the compressibility factor of a gas is using the CNGA equation as follows:
(Eq. 1-18)
Where
Pavg = Gas pressure, psig. [psig = (psia - 14.7)]
Tf = Gas temperature, 0R
G = Gas gravity (air = 1.00)
The CNGA equation for compressibility factor is valid when the average gas pressure Pavg is greater than 100 psig. For pressures less than 100 psig, compressibility factor is taken as 1.00. It must be noted that the pressure used in the CNGA equation is the gauge pressure, not the absolute pressure.
Example 1-8: Calculate the compressibility factor of a sample of natural gas (gravity = 0.6) at 80 0F and 1000 psig using the CNGA equation.
Solution:
From the Eq. (1.18), the compressibility factor is
The CNGA method of calculating the compressibility, though approximate, is accurate enough for most gas pipeline hydraulics work and process calculations.
1.9.2 Gas density at any condition of Pressure and temperature
Once the molecular weight of the gas is known, the density of a gas at any condition of temperature and pressure is given as:
ρg= ((MW)P)/RTZ lb/ft3
Since R=10.73, then
ρg= 0.093 ((MW)P)/TZ lb/ft3 (Eq. 1-19)
where
ρg = density of gas, lb/ft3,
P =pressure, psia,
T =temperature, 0R,
Z =gas compressibility factor,
MW=gas molecular weight.
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Basics of Gas Field Processing Book
Prediction and Inhibition of Gas Hydrates Book
Basics of Corrosion in Oil and Gas Industry Book
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Chapter 1 11
Basics of Oil and Gas Treatment 11
1.1 Introduction 11
1.2 Hydrocarbon preparation 11
1.3 Produced Hydrocarbon Fluids 12
1.3.1: Hydrocarbon gases 14
1.3.2: Molecular weight and apparent molecular weight 16
1.3.3: Apparent molecular weight of gas mixture 17
1.3.4: Gas Specific Gravity and Density 18
1.3.5: General Gas Law 19
1.4 Natural Gas Field Processing 19
1.5 Natural Gas Composition 21
1.6 The heating value of gases 23
1.7 Natural Gas Sampling 24
1.7.1 General Overview and introduction 24
1.7.2 Sample procedures and precautions 26
1.8 Product specifications 32
1.8.1 Natural gas 32
1.8.2 Natural-Gas Liquids 35
1.9 Physical properties of Hydrocarbon Gases 35
1.9.1 Compressibility Factor (z) 35
1.9.2 Gas density at any condition of Pressure and temperature 38
1.9.3 Gas volume at any condition of Pressure and temperature 39
1.9.4 Velocity of gas, (ft/s) 41
1.9.5 Average pipeline pressure 42
1.9.6 Viscosity of gases 43
--------------------
Chapter 1
Basics of Oil and Gas Treatment
1.1 Introduction
Oil and gas wells produce a mixture of hydrocarbon gas, condensate or oil, salt water, other gases, including nitrogen, carbon dioxide (CO2), and possibly hydrogen sulfide (H2S), and solids, including sand from the reservoir, dirt, scale, and corrosion products from the tubing.
These mixtures are very difficult to handle, meter, or transport. In addition to the difficulty, it is also unsafe and uneconomical to ship or to transport these mixtures to refineries and gas plants for processing. Further, hydrocarbon shipping tankers, oil refineries, and gas plants require certain specifications for the fluids that each receive. Also, environmental constraints exist for the safe and acceptable handling of hydrocarbon fluids and disposal of produced salt water. It is therefore necessary to process the produced fluids in the field to yield products that meet the specifications set by the customer and are safe to handle.
1.2 Hydrocarbon preparation
The goal is to produce oil that meets the purchaser’s specifications that define the maximum allowable amounts of water, salt, and sulfur. In addition to the maximum allowable value of Reid vapor pressure and maximum allowable pour point temperature.
Similarly, the gas must be processed to meet purchaser’s water vapor maximum allowable content (Water dew point), hydrocarbon dew point specifications to limit condensation during transportation, in addition to the maximum allowable content of CO2, H2S, O2, Total Sulfur, Mercaptan, Mercury, and maximum gross heating value.
The produced water must meet the regulatory requirements for disposal in the ocean if the wells are offshore, or to meet reservoir requirements for injection into an underground reservoir to avoid plugging the reservoir.
The specifications for the above requirements may include maximum oil in water content, total suspended solids to avoid formation plugging, bacteria counts, toxicity in case of offshore disposal, and oxygen content. Before discussing the industry or the technology of oil and gas processing it is best to define the characteristic, physical properties and main chemical composition of oil and gas produced.
Oil field gas processing which is the subject of this book; generally consists of two distinct categories of operations:
Separation of the gas-oil-brine well-stream into its individual phases,
Removal of impurities from the separated phases to meet sales/transportation/reinjection specifications and/or environmental regulations.
Figures 1-1 and 1-2, illustrates gas-oil separation plant, and gas flow diagram.
Fig.1-1 .Gas Oil Separation plant function and gas flow diagram.
1.3 Produced Hydrocarbon Fluids
The desirable constituents of crude oil and natural gas are hydrocarbons. These compounds range from methane (CH4) at a lower-molecular-weight end all the way up to paraffin hydrocarbons with 33 carbon atoms and poly-nuclear aromatic hydrocarbons with 20 or more carbon atoms. Natural gas is principally methane. Crude oil is principally liquid hydrocarbon having 4 and more carbon atoms.
There are a tendency to regard crude oil as a liquid and natural gas as a gas and consider production of the two phases as separate operations. However, in the reservoir, crude oil almost always contains dissolved methane and other light hydrocarbons that are released as gas when the pressure on the oil is reduced. As the gas evolves, the remaining crude-oil liquid volume decreases; this phenomenon is known as shrinkage. The gas so produced is called associated or separator gas. Shrinkage is expressed in terms of barrels of stock-tank oil per barrel of reservoir fluid. Crude oil shrinkage is the reciprocal of oil formation volume factor (FVF).
Similarly, natural gas produced from a gas reservoir may contain small amounts of heavier hydrocarbons that are separated as a liquid called condensate. Natural gas containing condensate is said to be wet. Conversely, if no condensate forms when the gas is produced to the surface, the gas is called dry.
A hydrocarbon constituent range or a spectrum of well fluids usually produced are summarized in table 1-1 (as noted by McCain (1973)). The type of fluid produced depends on the phase diagram of the reservoir fluid and the reservoir temperature and pressure, as will be discussed later in phase behavior of natural gas.
Notes.
Separator may be a slug catcher, free water knock out drum, two phase separator, or gun barrel.
Oil and water are separated and undergoes further treatment processes not in the scope of this book.
Fig.1-2. An example of gas flow diagram
Figure 1-3 depicts a typical gas-oil separation sequence (including incidental water removal). Table 1-1 lists the five common types of wellstream fluids and summarizes typical yields and liquid properties.
Table. 1-1. Petroleum fluid spectrum (after McCain, 1990)
When crude oil is separated from its associated gas during production, the total gas evolved while reducing the oil to atmospheric pressure divided by the volume of the remaining crude oil is called gas-oil ratio or GOR. The GOR is expressed as the total standard cubic feet of gas evolved per 60 0F barrel of stock-tank or atmospheric-pressure oil (scf/bsto). SI units are standard cubic meters of gas per cubic meter of 15 0C oil (metric units).
The total GOR depends on the number of stages used in the separation sequence, as well as the operating pressure of each stage. For three or more stages, the GOR approaches a limiting value. Optimization of the separation sequence usually involves either maximizing crude-oil yield or minimizing recompression horsepower as well be explained briefly in chapter 3.
For wet natural gas, the liquid content is given in barrels of condensate per million standard cubic feet of gas (bbl/MMscf) or in U.S. gallons of condensate per thousand standard cubic feet (GPM).
The various types of produced hydrocarbons have been described using GOR (McCain, 1973- and after McCain, 1990. Table 1-1) or composition (Gould and McDonald, 1979 – Table 1-2).
Of course the gas yield depends on the relative amounts of the various hydrocarbons present, and perhaps Penick (1983) summarizes these relationships best in figure 1-3 ( originally drawn by W.H. Speaker, Jr.).
Allen (1952) emphasizes that classifications like Table 1-1. are an oversimplification in the sense that GOR does not always reflect the condition of the fluid in the reservoir. For example, in the GOR range of 3000 – 7000 scf/bbl the reservoir fluid may be a liquid or a denese fluid (highly-compressed gas) but still be a black oil on the surface. If the fluid is a denese fluid at reservoir temperature and pressure, then any liquid produced is usually defined as condensate. If the reservoir fluid is a liquid, then the produced liquid is called volatile oil. The difference in these two cases is explained in chapter 2.
Table -1-2 . Well-Fluid Type (Gould and McDonald, 1979)
1.3.1: Hydrocarbon gases
Most of compounds in crude oil and natural gas consist of molecules made up of hydrogen and carbon, therefore these types of compounds are called hydrocarbon.
The smallest hydrocarbon molecule is Methane (CH4) which consists of one atom of Carbon and four atoms of hydrogen. It may be abbreviated as C1 since it consisted from only one carbon atom. Next compound is Ethane (C2H6) abbreviated as C2, and so on Propane (C3H8), Butane (C4H10)...etc.
Hydrocarbon gases are C1:C4), with the increase of carbon number, liquid volatile hydrocarbon is found (e.g. Pentane C5 is the first liquid hydrocarbon at standard conditions).
Fig. 1-3. Typical Reservoir composition (Penick, 1983).
Fig.1-4. Stage separation (Gas-oil separator train).
1.3.2: Molecular weight and apparent molecular weight
The molecular weight of a compound is the sum of the atomic weight of the various atoms making up that compound. The Mole is the unit of measurements for the amount of substance, the number of moles is defined as follows:
Mole = Weight/(Molecular weight) (Eq. 1-1)
Expressed as n = m/M (Eq. 1-2)
or, in units as, lb-mole = lb/(lb/lb-mole) (Eq. 1-3)
Table 1-3 Physical constants of light hydrocarbons and some inorganic gases. Adapted from GPSA, Engineering Data Book, chapter 23.
Example 1.1: Methane molecule consists of one carbon atom with atomic weight =12 and 4 hydrogen atoms with atomic weight = 1 each. Molecular weight for Methane (CH4) = (1 × 12) + (4 × 1) = 16 lb/lb-mole. Similarly, Ethane (C2H6) molecular weight = (2 × 12) + (6 × 1) = 30 lb/lb-mole.
Hydrocarbon up to four carbon atoms are gases at room temperature and atmospheric pressure. Reducing the gas temperature and/or increasing the pressure will condense the hydrocarbon gas to a liquid phase. By the increase of carbon atoms in hydrocarbon molecules, consequently the increase in molecular weight, the boiling point increases and a solid hydrocarbon is found at high molecular weight.
Physical constants of light hydrocarbon and some inorganic gases are listed in Table 1-3.
1.3.3: Apparent molecular weight of gas mixture
For compounds, the term molecular weight is used, while, for hydrocarbon mixture the term apparent molecular weight is commonly used. Apparent molecular weight is defined as the sum of the products of the mole fractions of each component times the molecular weight of that component. As shown in Eq. 1-4
MW= ∑▒ Yi (MW)i (Eq. 1-4)
where
yi =molecular fraction of ith component,
MWi =molecular weight of ith component,
Ʃyi =1.
Example 1.2: Determine the apparent molecular weight for the gas mixture in Table 1-4:
Table 1-4 Gas mixture for Example 1-2
Solution: Using Table 1-3 & Equation 1-4
MW= ∑▒ Yi (MW)i
MW = (Mole Fraction of component 1 × MW of component 1) + (Mole Fraction of component 2 × MW of component 2) + (Mole Fraction of component 3 × MW of component 3) +…etc.
The following table can be made:
Table 1-5 Solution of Example 1-2
The apparent molecular weight is 21.36
1.3.4: Gas Specific Gravity and Density
The density of a gas is defined as the mass per unit volume as follows
Density = mass / volume (Eq. 1-5)
The specific gravity of a gas is the ratio of the density of the gas to the density of air at standard conditions of temperature and pressure.
S = (ρ(gas))/(ρ(air)) (Eq. 1-6)
Where
ρ(gas) ρg = density of gas
ρ(air) ρair = density of air
Both densities must be computed at the same pressure and temperature, usually at standard conditions.
It may be related to the molecular weight by Equation 1-7
S = (MW(gas))/(MW(air)) (Eq. 1-7)
Since molecular weight of air is 28.96 (table 1-3)
Specific gravity of gas S = (MW(gas))/28.96 (Eq. 1-8)
Example 1-3: Determine the specific gravity of the gas mixture in example 1-2.
Solution:
Apparent molecular weight of gas mixture is 21.36
Gas specific gravity = 21.36/28.96 = 0.7376
Since the gas is a compressible fluid, its density varies with temperature and pressure, calculating the gas density at a certain pressure and temperature will be explained after discussing the general gas law and gas compressibility factor.
1.3.5: General Gas Law
The general (Ideal) Gas equation, or the Perfect Gas Equation, is stated as follows:
PV = nRT (Eq. 1-9)
Where
P = gas pressure, psia
V = gas volume, ft3
n = number of lb moles of gas (mass/molecular weight)
R = universal gas constant, psia ft3/lb mole OR
T = gas temperature, OR (OR = 460 + OF)
The universal gas constant R is equal to 10.73 psia ft3/lb mole OR in field units.
Equation (1-9) is valid up to pressures of about 60 psia. As pressure increases above this level, its accuracy becomes less and the system should be considered a non-ideal gas equation of state.
PV = znRT (Eq. 1-10)
Where
z = gas compressibility factor.
1.4 Natural Gas Field Processing
The main constituent of natural gas is methane, desirable as a primary fuel. Sales gas also contains smaller amounts of the heavier hydrocarbons listed in table 1- 6. Often a portion of the heavier hydrocarbon can be recovered profitably in a field-gas processing plant as one or more liquid products. These liquefilable components ( or condensate) may be recovered as a single liquid stream that is transported to a separate plant for fractionation into stable products. Alternatively, in very large field units, fractionation is performed in the field. Common natural gas liquid (NGL) products are summarized below.
Table. 1- 6. Common natural gas liquid (NGL) products nad uses.
The permanet gases occuring in natural gas include nitrogen, helium, argon, hydrogen, and oxygen. Most natural gases contains some nitrogen, and a few have 30 mole percent or more. Nitrogen lowers the heat of combustion of the gas. Since natural gas is usually sold on the basis of energy content with a fixed minimum heating value, the nitrogen content is limited to fairly low amounts in commercially salable gas. The removal of nitrogen require expensive cryogenic processing, so too high a nitrogen content may render a gas unusable.
Many natural gases contain a few hundredth of a percent of helium. Helium has no deleterious effect other than lack of heating value. Separated helium is very valuable and, in the U.S. for example, suffficient lawsuits have awarded royalty payments for helium that could have been recovered economically from sales gases. Analysis also reported occasional small amounts of oxygen, as well as argon and hydrogen. Chromatographic analyses may lump all the inert gases as nitrogen.
Hydrogen sulfide and carbon dioxide are found in many Natural gases and may occur in very high percentages. In fact, essentially pure carbon dioxide is produced, dehydrated, and pipelined for CO2 floods in the enhanced recovery of crude oil. Hydrogen sulfide and carbon dioxide are referred to as acid gases because they dissociate upon solution in water to form acidic solutions. Hydrogen sulfide is very toxic and corrosive, while carbon dioxide is corrosive.
A natural gas containing no hydrogen sulfide is said to be sweet. Conversely, a sour gas containing hydrogen sulfide. Strictly speaking, “sweet” and “sour” refer to both acid gases (CO2 and H2S) but are usually applied to hydrogen sulfide alone.
Sulfur compounds, other than H2S, are present in minute amounts and can affect field processing. These compounds tend to concentrate in the condensate and sometimes require treating (or sweetening) of the liquid products.
Removal of hydrogen sulfide to very low content (4 ppmv or 1/4 grain/100 scf) is required in the field. Carbon dioxide can be tolerated to much higher levels, say 1-2%, as long as the heating value of the sales gas is satisfactory.
There are many so-called “treating” processes for sweetening natural gas. These processes are either batch, reactant-discarded processes for removing low amounts of hydrogen sulfide or continuous solvent-regenerated processes for large amounts. Batch processes are used when the consumable-chemical cost is not prohibitive. The principal continuous, solvent regenerated treating processes use water solution of chemical solvents, typically alkanolamines. Other sweetening such as physical solvents, mixed physical-chemical solvents, or membranes may be more economical in some cases.
Because these processes also remove appreciable amounts of other sulfur compounds and/or carbon dioxide, many gas streams need no further processing to meet total sulfur and acid gas specifications.
Hydrogen sulfide is an extremely toxic substance. Fortunately, the familiar sulfurous smell can be detected at concentration less than 1 ppmv. However, extended exposure at higher levels deadens the sense of smell, so that odor alone cannot be a reliable detector.
Water or brine is undoubtedly present in many gas reservoirs but usually is not entrained to the surface. If free liquid water or brine is produced, a wellhead knockout drum vessel, or separator, is needed to prevent the water entering the gathering lines. At high pressure and low temperature, natural gas and free liquid water form solid hydrates capable of plugging flow lines. Produced water also may contain methanol and/or corrosion inhibitors that have been injected in the well string.
Produced gas is regarded as being saturated with water vapor at the wellhead conditions of temperature and pressure, even if no liquid water is produced. Associated gas is regarded as being saturated with water vapor at the outlet from the gas-oil separator in which it is produced. Water condensed downstream of the wellhead or gas-oil separator will be essentially pure (fresh) water rather than saline brine.
Water vapor is a contaminant that must be removed by proper processing of the natural gas stream. Gas transmission lines often specify water content of 7 lb/MMscf (or a water dew point of 32 0F or less). Triethylene glycol (TEG) is almost always used for such applications. Cryogenic plants require “bone-dry” gas (water dew point as low as -150 0F). Solid desiccants dehydration is typically used for dew points below – 25 0F.
Formation solids are not produced with most natural gas. Nevertheless, solids are sometimes separated from the process plant inlet separator. Usually these solids are mainly mill scale and rust from pipe wall, along with iron sulfide (for sour gas). The solids interfere with treating and dehydration processes and should be removed in a scrubber, filter, or possibly a filter-coalescer separator.
Very deep, high temperature, high pressure sour gas, may contain solid sulfur. Sulfur is inert but must be properly separated to prevent downstream processing problems.
Mercury has been detected in natural gas streams, in concentrations from approximately 1 ppbw to 230 ppmw. Mercury usually causes no processing problems but has caused corrosion of aluminum tubes in heat exchangers. Removal using sulfur-impregnated activated carbon, sulfur beds, or molecular sieves has been suggested.
Recently, arsenic compounds were reported present in natural gas produced from some fields. The discovery was made indirectly by the detection of white powder on regulators in gas company. The arsenic will be removed using a vertical copper-zinc adsorbent bed.
Methanol (to prevent hydrate formation) and corrosion inhibitors are sometimes added downhole and so may be present in natural gas at the wellhead. Any compressed gas will contain some entrained lubricating oils. Field processing may also introduce additional contaminants such as glycols and amines. Mixtures of these liquids with the previously listed solid contaminants are called sludge.
1.5 Natural Gas Composition
The gas analyses shown in table 1-7, span the composition ranges normally encountered. These analyses are typical of the data furnished to the designer of surface processing equipment. As previously mentioned, the heavier hydrocarbon in these gases are regarded as a recoverable liquids. The amount of potentially recoverable liquid is expressed as gallons liquid at 60 0F, if totally condensed, per 1000 standard cubic feet of gas (so called GPM, not to be confused with gallons per minute). A gas is termed lean or rich as follows:
Lean < 2.5 GPM
Moderately-Rich 2.5-5 GPM
Very Rich > 5 GPM
The above classification is based on ethane and heavier hydrocarbons (C2+) because ethane is sometimes regarded as a desirable feed for petrochemical process and can be recovered as a liquid in expander-type gas processing plants. If ethane is not considered as a valuable liquid component, the GPM can be based on propane and heavier hydrocarbons (C3+).
Table. 1-7. Typical gas analyses (Mol Percent).
Example 1-4. Confirm the GPM given in table 1-7. For the first natural gas stream.
Solution: First find the scf/gal of the C7+ components
scf C7+/ gal liq C7+ = (SG C7+)*(lb H2O/gal)*(1/Mol.Wt.)*(scf/lb mol) (Eq. 1-11)
= (0.803) * (8.334) * (1/172) (379.5) = 14.77
(379.5 = n and is a constant of (scf/lb mol) derived from PV=nRT at standard condition)
The table that follows details the calculations (which are most conveniently done with a spread- sheet program)
Table. 1-8. Solution of example 1-4.
(a) From Table 1-3.
(b) Column 3/ column 4
GPM is approximately = 0.3 * [100 – (C1 + N2+CO2)].
The weak link in gas analysis is often the composition of the C6+ or C7+ portion of the gas. For many purposes the small amount of the C6+ material renders its characterization unimportant. One important exception is when a full wellstream gas is to be transported in a pipeline over a long distance, such as from an offshore platform. Condensation of liquids in the line will cause a large pressure drop that must be anticipated if adequate platform compression is to be furnished and the proper pipe diameter selected. Very accurate characterization of the C6+ is needed for hydrocarbon (HC) dew points prediction, otherwise the HC can be determined in laboratory.
1.6 The heating value of gases
The heating value of a gas is expressed in Btu/ft3. It represents the quantity of heat in Btu (British Thermal Unit) generated by the complete combustion of one cubic foot of the gas with air at constant pressure (1 atmosphere = 14.7 psia) and at a fixed temperature of 60 0F.
Hydrogen in the fuel burns to water and when the flue gases are cooled to 60°F, the physical state — either vapor or liquid — of this water must be assumed. So the latent heat of vaporization of the water may or may not be considered to be part of the heating value. The result is two definitions for the heating value. The higher or gross heating value, HHV, includes the heat of condensation and the lower or net heating value, LHV, assumes the water remains in the vapor state.
For gas mixture the heating value is calculated as follows:
H = Ʃ xi Hi Eq. 1-12
Example 1-5: Calculate the heating value of gas mixture of Example 1-2
Table 1-9 Solution of Example 1-5
From table 1-9 the Gross calorific value HHV = 1246 Btu/ft3
The higher, ideal, dry heating value of sweet natural gas at 60°F and 760 mm Hg may be calculated with the following equation:
HHV=1568.72 × SG – 2524.88 × XCO2 – 1658.37 × XN2 +141.05 (Eq. 1- 13)
Applying Eq 1-13 for Example 1-5
The apparent molecular weight= 21.36
Gas Specific gravity = 21.36/28.96 = 0.738
HHV = 1568.72 × 0.738 – 2524.88 × 0.015 – 1658.37 × 0.01 +141.05 = 1244 Btu/ft3
1.7 Natural Gas Sampling
1.7.1 General Overview and introduction
The purpose of natural gas sampling is to secure a representative sample of the flowing gas stream for a specific period of time. Naturally, the more often the sampling system samples the flowing stream, the more likely it is to be truly representative of a stream with varying composition. Refer to figure 1-5.
Sampling systems consist of numerous components and must include some key elements including a sample probe, any necessary connecting tubing, sample containers or sample valves and appropriate heat-tracing and insulation. For online gas chromatographs, or on-stream analyzers, the sampling system ends at the injection valve on the inlet of the g.c.
The sample may be collected on a spot, composite or continuous basis.
Spot sampling simply means that a technician manually collects a sample directly from the stream at scheduled intervals or as needed.
Composite samples are usually collected on a weekly or monthly basis. Composite sampling systems should grab small samples on a flow proportional basis, then inject them into the composite sampling cylinder. There are composite sampling systems that work on a simple time cycle (time proportional sampling), but they are not recommended, especially if they continue to sample even when flow has stopped. If a time proportional system is already in service, it must be equipped with a flow switch or similar device to ensure that when flow stops, sampling will stop.
Continuous systems provide a steady flow of sample through a sample loop that passes near a composite sampler or on stream analyzer. In the case of an on-line chromatograph, the injection valve of the g.c. is able to admit and distribute a sample from the flowing loop that is representative of the flow in the main line. Sample rate flow loops must be carefully sized and generally should operate at velocities around 5 ft/sec., but this may vary if the sample loop is exceptionally long (over 100 feet).
The sample point is usually located downstream of the meter run and must be remote from severe flow disturbances such as control valves and orifice plates by at least five nominal pipe diameters. For flowing streams that are not near their hydrocarbon dew point, the probe should be positioned either upstream or downstream of the meter tube, and at least 5 pipe diameters downstream of any flow disturbing elements, such as elbows, swirl generators, headers, valves and tees. If the sample source is at or near its hydrocarbon dewpoint, some research has indicated that the probe should be located at least 8 pipe diameters downstream of any flow disturbance, including an orifice meter. The sample point must not be installed within the upstream or downstream engineered sections of the meter tube, since the fitting and probe could produce disturbances in the flow profile going through the meter in the run. The probe installed in the sample point extends into the center 1/3 of the internal diameter of the meter run to ensure no heavy materials or contaminants migrating along the pipe walls are allowed to contaminate the sample. Note that for large diameter pipelines, the probe never needs to be longer than 10 inches. The probe is equipped with an outlet valve to allow the system to be shut in when no sampling is being performed or to perform maintenance on downstream equipment in composite or continuous sampling systems.
The tubing connecting the sample probe to the downstream sample system(s) should be internally clean, as short as practical (usually 6 to 24 inches maximum) and made of either nylon or stainless steel. Stainless steel is actually preferred due to its strength and flexibility and resistance to melting and/or sharp edges, but nylon is not porous and when used safely, can also give good analytical results. Teflon, carbon steel, plastic tubing and many other materials do not perform well.
Care must be taken to insure there are no leaks in the sampling system. Typically, if a leak occurs, smaller molecules tend to escape preferentially and create a bias in analytical results. If the leak is large, there may be enough cooling to produce condensation in the sample system and cause the samples to be very non-representative.
Note that whenever the sample line is operating in ambient temperatures below the flowing temperature of the stream, the line may need to be heat-traced and insulated. If the ambient temperature is lower than the dewpoint temperature of the flowing stream, heat tracing and insulation are required.
Be sure that the heat tracing is properly and safely done, using electrically limited tracing meeting appropriate electrical codes for the area classification (typically Class I Group D Division I or II).
Realize that the dewpoint of a gas stream is absolutely critical to accurate sampling. If any component in the sampling system causes the temperature of even a portion of the gas stream being sampled to cool to or below the hydrocarbon dewpoint, the sample will be depleted of heavy components and can no longer be truly representative of the stream. Note that the Btu content in this situation is not always too low. If the sample system continues to condense heavy components for an extended amount of time, accumulations may reach the point that liquid droplets enter the sample and actually cause the indicted Btu content and calculated relative density (specific gravity) to be too high.
The sample cylinders used in spot sampling should be stainless steel, single cavity cylinders. Single cavity cylinders are recommended, due to the difficulty of fully cleaning piston cylinders between uses. Residue that may remain in the piston cylinders and their seals may produce incorrect analyses. The cylinders should be equipped with standard design sample valves that are screw open or closed (not 1/4 turn ball valves) and have a flow passage of approximately 1/8 inch diameter.
It may be convenient to use piston (constant pressure) cylinders in composite sampling systems, since you can easily see that the system is working or not working as the level indicator moves. If constant pressure/piston cylinders are used and oil or grease contamination is present in the system, they must be disassembled between uses, carefully cleaned and then the seal rings must be replaced if the cylinder is expected to provide representative samples.
Note that if sampling is being performed to determine the levels of volatile or reactive contaminants, such as H2S, the cylinder may need to be lined with an epoxy/phenolic lining.
Even then, particularly reactive materials, such as H2S or ethyl mercaptan are likely to be lost prior to analysis unless the sample is collected on-site and analyzed immediately.
Even a few minutes delay can reduce detectable levels of reactive materials. Shipping a sample to a remote lab and delaying analysis beyond a couple of hours will essentially ensure that the indicated levels of the reactive/volatile material will be too low or perhaps not detectable at all.
The two spot sampling methods that are most recommended are the fill and empty method and the helium pop method. The displacement methods also performed reasonably well during the recent API research studies.
The fill and empty method requires that the cylinder be equipped with a “pigtail” following the sample cylinder outlet valve. While leaving the sample cylinder inlet and outlet valves open, the probe outlet valve and the valve at the end of the pigtail are cycled to alternately fill and empty the sample cylinder. The pigtail ensures that the heat of compression created when the sample cylinder is filled more than offsets the Joule-Thomson cooling produced when the sample cylinder is depressurized. It does this by insuring the maximum pressure drop while depressurizing the sample cylinder is far removed from cylinder itself, at the end of the pigtail. The ability of the fill and empty procedure to actually elevate the temperature of the sampling cylinder above the flowing temperature of the stream being sampled during many operating conditions makes this method the most desirable when the ambient temperature is at or near the hydrocarbon dewpoint of the stream.
The pigtail should be approximately 1/4 inch tubing and be at least 36 inches in length, although it may be coiled to make the apparatus easier to handle. The coils should not touch one another, otherwise the heat loss at the end of the pigtail may be transferred quickly across the coils to the sample cylinder.
There should be another sample valve, similar to the sample cylinder valves on the outlet of the pigtail. The flow passage through this valve must not be larger than the passage through the cylinder valves. Refer to API Chapter 14.1 for the detailed procedures for performing the fill and empty method and to either API Chapter 14.1 or GPA 2166 for the number of fill and empty purge cycles required at various line pressures.
1.7.2 Sample procedures and precautions
Collection of a truly representative samples is not a simple job as it looks like. A moment’s thought will confirm that any subsequent design, operating, or investment decision can be no better than the prior sampling and analyses.
In spite of its obvious importance, natural-gas sampling is seldom done well. Natural gas, flowing with accompanying condensate, can not be properly sampled by withdrawal of a portion through a sample tap. Reflection of the complex natural of two phase flow (with its varying flow patterns and liquid-holdup phenomenon, along with velocity gradients in the fluid) should convince one of the impossibility.
In the ideal case, a separator should be set up at the wellhead to collect condensate and meter each phase. Subsequent analysis of the separated gas and liquid allows recombination to obtain the wellstream analysis. This is the reason well-test separators are even installed offshore where space and weight are extremely expensive.
Poor sampling is generally caused by:
Ignorance of its economic and technical importance.
Inadequate training of personnel involved.
Ignorance and/or unwillingness to follow standard procedures.
While sampling in the field do not:
Use dirty cylinders that contain previous samples.
Sample when pressure and temperature are not stable
Sample when separator pressure is different from sample pressure.
Dilute sample with air.
Fail to identify sample completely.
Sampling hydrocarbon fluids can be hazardous. Every body involved in sampling must be familiar with and follow safe practices while handling flammable fluids under pressure. All sample containers must meet the specification to handle sampled fluids, and must be clearly labeled.
Do not fill liquid sample containers completely full. A 150 ml, stainless-steel container was filled full with gasoline and closed at 0 psig and 61 0F. When heated to 105 0F, the sample pressure was 3100 psig!.
1.7.2.1 Gas sampling
GPA standard 2166-86 indicates that good samples can be obtained by all 8 approved methods provided extreme care is taken, however the sampling project report ranked the accuracy of the eight methods in the following order of decreasing accuracy (refer to fig. 1-5):
Water displacement
Continuous purge
Purge and fill
Glycol displacement
Reduced pressure
Floating piston cylinder
Helium filled to 5 psig
Evacuated container
In practice purge and fill is the most popular method while the floating piston cylinder is finding increased use because composite samples cam be obtained.
Several problems can arise during sampling of natural gas:
Condensation of hydrocarbons due to temperature and pressure changes during sampling.
Entrainment of liquid droplets and mist
Sample constituents can react with sample container.
Some sample components may dissolve in the displacement media.
Subsequent handling of the sample is therefore, very important to assure the relatively small amounts of condensable components remain in the gas phase and are removed to the chromatograph or mass spectrometer.
In a natural gas, hydrogen sulfide in small amounts may be overlooked entirely, creating a future problem in design and operation. Hydrogen sulfide reacts rapidly with carbon steel and may disappear from the sample altogether. A stainless-steel sample cylinder should be used if the presence of H2S is suspected. Even austenitic stainless-steel will absorb a small amount of H2S. GPA Standard 2261 recommends that the gas to be analyzed at its source for hydrogen sulfide content less than 3% by GPA Standard method 2377. Teflon-lined cylinders have also been used successfully.
The following precautions are recommended:
Use a pre-filter and/or “knockout” trap on the sample system just downstream from the source valve. This is mandatory on “wet” gas samples.
Use heated sample lines to prevent condensation from low ambient temperature during sampling.
Keep sample lines as short as possible. This should be done for all sample grabbing, wet or dry.
Clean, dry, and evacuate sample cylinders before taking to the field. This prevents possible liquid carryover from the previous sample.
Purge sample cylinder carefully.
Empty cylinders through a non-metallic pigtail on the outlet for the expansion of pressured gas to the atmosphere.
For samples requiring calorimetric heating value determination and chromatographic analysis, use stainless steel or carbon steel, DOT3A and DOT3AA, cylinders, 300 cu.in, volume. Smaller cylinders may be used for GC analysis only; however, a calorimeter Btu determination provides good confirmation of the chromatographic analysis. Calculated Btu from the analysis should not be over +/- 3 Btu from the measured value.
Keep sample cylinders upright while filling. With the purge line at the bottom. Do not lay the cylinders on the ground.
Use line probe that extends at least a third of the way into the line of sampling. This prevents the entry condensate and other contaminants into the sampling line.
Sampling by liquid displacement has been used in the past and still followed by some organizations. It is not recommended for gases or liquids containing acid-gases such as CO2 and H2S. These components dissolve in the liquid, usually water, and are lost. Therefore the resulting analysis is not representative.
Continuous sampling
The following procedures and precautions are recommended for continuous or composite sampling of natural gas:
Sample point should ”see” center one-third of the pipeline in an area of good velocity with minimum turbulence.
Sample probe, equipped with full open ball or gate valve, must be kept away from pipeline fittings and orifice plates. The probe may be bevel or flat cut at end, but must be kept clear of free liquids and aerosols. Bevel may face upstream or downstream.
Probe construction: stationary or permanent, manually insertable or automatic insertion type probe s may be used. The probe should be stainless steel so that it will not react with the sampled gas.
Sampler hook-up and manifold: The sampler should be mounted above the sample point as shown in figure 1- 6. Other precautions are:
a. Line from probe to sampler should slope down to let any free liquid to drain back into pipeline.
b. Never sample a dead-end line.
c. Any leaks in the line from the sampler to the cylinder will “lose” light ends preferentially.
d. Any filters, drips, or regulators between the probe and sampler will invalidate the sample.
Sampler: Should take a composite sample in the same way as an operator would take a spot sample. If the pipeline flow rate varies, the sampler should be actuated proportional to the flow. The sampler should be able to pump the sample into the cylinder and purge itself before pumping each new “bite”.
Sample containers
Proper cleaning and inspection of sample cylinders cannot be overemphasized. Several methods are available:
Volatile solvent and air dry
Steam clean and air dry
Evacuation
Evacuation and fill with 5 psig helium
Correct procedures must be followed rigorously in every detail.
FIG. 1-5. Fill and Empty Sampling Method, and helium pop sampling method.
Fig.1-6 .Gas automatic or continuous sampler.
1.7.2.2 Natural Gas Liquid Sampling
Utilizing Floating Piston Cylinders
Liquid sampling requires special precautions to accumulate and transfer representative samples. Pressure in the sample cylinder and/or accumulator must be maintained at 1.5 times the product vapor pressure. Maximum product vapor pressure should be determined using the highest ambient temperature or flowing temperature (whichever is highest) to determine the minimum pre-charge pressure. A method to break up stratification must be provided prior to transfer of the sample to another container and laboratory analysis. Maintaining the appropriate pressure and mixing the sample can be satisfied by using floating piston sample cylinders with mixers (Fig. 1-7). A “rattle ball” or agitator may be used in place of the mixing rod shown. The floating piston cylinders are precharged on one end with an inert gas at a pressure 1.5 times above product vapor pressure. This prevents sample vaporization, which could result in erroneous analysis. This design also provides a compressible inert gas cushion to allow for thermal expansion of the liquid. A pressure relief valve is needed, but should it discharge, the integrity of the sample will be lost.
Liquid sample cylinders shall not be filled over 80 percent full.
FIG. 1-7. Liquid Sample Container
Samples are acquired through a sample probe inserted into the center third of the flowing stream. The probe should be mounted in the top or side of the line. Continuous samplers should use a continuous flowing sample loop (speed loop) or a probe mounted sample pump to ensure the most current sample is always added to the sample container. Speed loops must have a driving device such as an orifice, differential pump, available pressure drop, or “scoop” probes. The driving device should be sized to provide a complete exchange of liquid in the sample loop once per minute. The sample pump must be set to gather flow proportional samples to ensure a true representative sample is obtained. If flowing pressures are higher than the sample accumulator pre-charge pressure, then the sample pump must prevent “free-flowing” of product into the sample container. See Fig. 1-8 for an example continuous sampling application. Speed loop lines may require insulation when cold ambient temperatures have a significant effect on viscosity. The product in the sample container must be thoroughly mixed before being transferred to a transport cylinder.
Details and alternative methods for obtaining liquid samples are found in GPA 2174.
Fig. 1-8. Continuous Sampler (Automatic sampler)
Procedures and precautions for liquid sampling
A GPA Work group conducted a cooperative natural gas liquid sampling project for future revision of GPA Standard 2174. The following four sampling methods were judged acceptable:
Floating Piston Cylinder.
Water displacement (total removal – 80% hydrocarbons/20%displaced outage)
Water displacement (partial removal – 70% hydrocarbons/20%displaced outage/10%water remaining in cylinder)
Ethylene glycol displacement (total removal – 80% hydrocarbons/20%displaced outage)
The following precautions are recommended for liquid sampling while using the floating piston cylinder:
Pressure on the backside piston should be higher than the line pressure at the beginning of operation
Slowly bleed pressure down to enter sample, and maintain pressures on each side of the piston at nearly equal levels with just enough difference to move the pistons
Do not bleed the pressure off the backside of the piston after sample has been taken, this will flash some of the liquid into gas. Then you will need to resample
Do not fill cylinder 100% full with sample. Leave at least 25% as a pressure buffer to take care of ambient temperature fluctuations.
The floating piston cylinder can be used for “wet” gas sampling with excellent results in getting representative samples. Procedures are as follows:
Pressure up back of cylinder with line gas to full line pressure.
Connect cylinder to source (use of a line probe is mandatory) and open valves to cylinder.
Fill cylinders as you would in taking liquid sample. Note: It is extremely important to avoid any pressure drops. The pressure differences should not exceed 2 psi maximum. This will prevent flashing of heavy components which may be in aerosol form. Sample should be slowly entered into cylinder.
1.8 Product specifications
The objective of field processing is to provide transportable and/or salable fluids. There are two main products: natural gas and condensate (raw mix) or natural gas liquid (NGL).
1.8.1 Natural gas
Table 1-10 lists typical natural gas pipeline specifications listed in a typical sales gas contract. Theses specifications are fixed by negotiation between seller and buyer and vary from case to case. Not all sales gases will have all the specifications shown in the following items.
Table. 1- 10. Natural gas pipeline specifications.
Table. 1-11. Is a typical sales gas/ pipeline contract between two middle east companies.
Table. 1-11. Typical sales gas/ pipeline contract.
Wobbe Index is some times used in sales gas specification. Wobbe Index = [Gross heating value / (Sp.Gr)0.5].
The most important specifications are: water content (water dewpoint), H2S content, and gross heating value. The following table summarizes the effect of water, acid gases, and liquid hydrocarbons in sales gases:
Table.1-13. Water dew point and water content relation. (ANSI/CGA G-7.1 -1989 document).
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Note that both water and hydrogen sulfide must be removed to very low concentrations. Heating value is more complex (specification usually from 950 to 1200 Btu/scf).
Refer to table 1-3. For heating value of hydrocarbon gases. (Complete physical properties of hydrocarbon gases are listed in GPSA, Engineering data book, chapter 23).
From table 1-3, we can realize that; the most abundant component, methane, has a relatively low heating value (1010 Btu/scf). Therefore, by itself methane cannot always fulfill the minimum heating value requirement when inert gases (nitrogen and/or carbon dioxide) are present. However, enough heavier hydrocarbons are usually present to provide the required heating value, even with condensate recovery.
When there is a maximum heating value or hydrocarbon dew point specifications, some of the heavier hydrocarbons constituents may have to be removed as condensate. Refrigeration to (-30 0F reduces the heating value) how much, of course, depends on the gas composition and pressure. Further reduction requires cryogenic processing. The remaining specifications are met by suitable processing. These processes are the subjects of the later chapters.
Fig. 1-9 Moisture Content Nomograph for Gases
1.8.2 Natural-Gas Liquids
Hydrocarbon condensate recovered from natural gas may be shipped without further processing or stabilized to produce a safely-transportable liquid. In the case of raw condensate, there are no particular specifications for the product other than the process requirements. Stabilized liquid, on the other hand, will generally have a vapor pressure specifications, since the product will be injected into a pipeline or transport pressure vessel which has a definite pressure limitations.
Natural-gas liquid products are prepared by fractionation of the raw make into appropriate products, either at the field processing site or, perhaps more commonly, at a large central facility. In any case, product specifications are not so typical as those of sales gas, but depends heavily on the particular contract. Liquid product specifications generally include composition, vapor pressure, water content, and sulfur content.
1.9 Physical properties of Hydrocarbon Gases
1.9.1 Compressibility Factor (z)
The Compressibility factor, Z is a dimensionless parameter less than 1.00 that represents the deviation of a real gas from an ideal gas. Hence it is also referred to as the gas deviation factor. At low pressures and temperatures Z is nearly equal to 1.00 whereas at higher pressures and temperatures it may range between 0.75 and 0.90. The actual value of Z at any temperature and pressure must be calculated taking into account the composition of the gas and its critical temperature and pressure. Several graphical and analytical methods are available to calculate Z. Among these, the Standing-Katz, and CNGA methods are quite popular. The critical temperature and the critical pressure of a gas are important parameters that affect the compressibility factor and are defined as follows.
The critical temperature of a pure gas is that temperature above which the gas cannot be compressed into a liquid, however much the pressure. The critical pressure is the minimum pressure required at the critical temperature of the gas to compress it into a liquid.
As an example, consider pure methane gas with a critical temperature of 343 0R and critical pressure of 666 psia (Table 1-3).
The reduced temperature of a gas is defined as the ratio of the gas temperature to its critical temperature, both being expressed in absolute units (0R). It is therefore a dimensionless number.
Similarly, the reduced pressure is a dimensionless number defined as the ratio of the absolute pressure of gas to its critical pressure.
Therefore we can state the following:
Tr = T/Tc (Eq. 1-14)
Pr = P/Pc (Eq. 1-15)
Where
P = pressure of gas, psia
T = temperature of gas, 0R
Tr = reduced temperature, dimensionless
Pr = reduced pressure, dimensionless
Tc = critical temperature, 0R
Pc = critical pressure, psia
Example1-6: Using the preceding equations, the reduced temperature and reduced pressure of a sample of methane gas at 70 0F and 1200 psia pressure can be calculated as follows
Tr = (70 +460) / 343 =1.5
Pr = 1200/666 = 1.8
For natural gas mixtures, the terms pseudo-critical temperature and pseudo-critical pressure are used. The calculation methodology will be explained shortly. Similarly we can calculate the pseudo-reduced temperature and pseudo-reduced pressure of a natural gas mixture, knowing its pseudo-critical temperature and pseudo-critical pressure.
The Standing-Katz chart, Fig. 1.10 can be used to determine the compressibility factor of a gas at any temperature and pressure, once the reduced pressure and temperature are calculated knowing the critical properties.
Pseudo-critical properties allow one to evaluate gas mixtures. Equations (1-16) and (1-17) can be used to calculate the pseudo-critical properties for gas mixtures:
P’c = Ʃ yi Pci (Eq. 1-16)
T’c = Ʃ yi Tci (Eq. 1-17)
where
P’c =pseudo-critical pressure,
T’c =pseudo-critical temperature,
Pci =critical pressure at component i, psia
Tci =critical temperature at component i, 0R
Yi =mole fraction of each component in the mixture,
Ʃ yi =1.
Example 1-7:
Calculate the Compressibility factor for the following Gas mixture at 1000F and 800 psig:
Table 1-14 for Example 1-7.
Using Equation 1-16 and 1-17
T`r = (100+460)/464.5 =1.2
P`r = (800+14.7)/659.4 = 1.23
From fig.1-10. Compressibility factor is approximately, z= 0.72
Figure 1-10 Compressibility Factor For lean sweet natural gas (Surface Production Operations).
(More graphs for compressibility factors and acid gas corrections are available in GPSA, data book)
Calculating the compressibility factor for example 1-6, of the gas at 70 0F and 1200 psia, using Standing-Katz chart, fig. 1-10. Z = 0.83 approximately. For ) Tr = 1.5 , Pr = 1.
.Another analytical method of calculating the compressibility factor of a gas is using the CNGA equation as follows:
(Eq. 1-18)
Where
Pavg = Gas pressure, psig. [psig = (psia - 14.7)]
Tf = Gas temperature, 0R
G = Gas gravity (air = 1.00)
The CNGA equation for compressibility factor is valid when the average gas pressure Pavg is greater than 100 psig. For pressures less than 100 psig, compressibility factor is taken as 1.00. It must be noted that the pressure used in the CNGA equation is the gauge pressure, not the absolute pressure.
Example 1-8: Calculate the compressibility factor of a sample of natural gas (gravity = 0.6) at 80 0F and 1000 psig using the CNGA equation.
Solution:
From the Eq. (1.18), the compressibility factor is
The CNGA method of calculating the compressibility, though approximate, is accurate enough for most gas pipeline hydraulics work and process calculations.
1.9.2 Gas density at any condition of Pressure and temperature
Once the molecular weight of the gas is known, the density of a gas at any condition of temperature and pressure is given as:
ρg= ((MW)P)/RTZ lb/ft3
Since R=10.73, then
ρg= 0.093 ((MW)P)/TZ lb/ft3 (Eq. 1-19)
where
ρg = density of gas, lb/ft3,
P =pressure, psia,
T =temperature, 0R,
Z =gas compressibility factor,
MW=gas molecular weight.
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Fundamentals of Oil and Gas Processing
Basics of Gas Field Processing
Basics of Corrosion in Oil and Gas Industry
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Basics of Gas Field Processing
Basics of Corrosion in Oil and Gas Industry
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Re: Basics of Gas Field Processing Book "Full text"
Chapter 1 - Part 2
Example 1-9: Calculate the pseudo-critical temperature and pressure for the natural gas stream composition given in example 1-2, calculate the compressibility factor, and gas density at 600 psia and 1000F.
Solution:
Table 1-15 solution of Example 1-9.
From the table MW= 21.36
T`c = 451.5 0R
P`c = 667 psia
From Eq. (1-14) and Eq. (1-15)
Tr = T/T`c = (100+460)/451.5 = 1.24
Pr = P/P`c = 600/667 = 0.9
Compressibility factor z could be calculated from figure 1-10, or from Eq. (1-18)
Value from figure, z = 0.83
From Equation 1-15 z = 0.87
For our further calculation we will use the calculated z value [Eq. (1-18)]
Using eq. (1-19) density of gas
ρg = 0.093 ((21.36)600)/(560 ×0.83) = 2.56 lb/ft3
Comparing ρg at standard condition (z=1)
ρg at standard condition = 0.093 (21.36)14.7/(520 ×1) = 0.056 lb/ft3
We can conclude that density increases with pressure while the volume decreases.
1.9.3 Gas volume at any condition of Pressure and temperature
Volume of a gas is the space occupied by the gas. Gases fill the container that houses the gas. The volume of a gas generally varies with temperature and pressure.
Volume of a gas is measured in cubic feet (ft3).
Gas volume are commonly referred to in "standard" or "normal" units.
Standard conditions commonly refers to gas volumes measured at: 60°F and 14.696 psia
The Gas Processors Association (GPA) SI standard molar volume conditions is 379.49 std ft3/lb mol at 60°F, 14.696 psia.
Therefore, each mole (n) contains about 379.5 cubic feet of gas (ft3)at standard conditions.
Therefore, by knowing the values of mass and density at certain pressure and temperature, the volume occupied by gas can be calculated.
Example 1-10: Calculate the volume of a 10 lb mass of gas (Gravity = 0.6) at 500 psig and 80 0F, assuming the compressibility factor as 0.895. The molecular weight of air may be taken as 29 and the base pressure is 14.7 psia.
Solution:
The molecular weight of the gas (MW) = 0.6 x 29 = 17.4
Pressure =500+14.7 = 514.7 psia
Temperature = 80+460 = 540 0R
Compressibility factor z= 0.895
The number of lb moles n is calculated using Eq. (1-2). n=m/(MW)
n = 10/17.4
Therefore n= 0.5747 lb mole
Using the real gas Eq. (1-10), PV=nzRT
(514.7) V = 0.895 x 0.5747 x 10.73 x 540. Therefore, V = 5.79 ft3
Example 1-11: Calculate the volume of 1 lb mole of the natural gas stream given in the previous example at 1200F and 1500 psia (compressibility factor Z = 0.811).
Solution:
Using Eq.(1-10), PV = nzRT
V= 0.811 x 1 x 10.73 x (120+460)/1500. V = 3.37 ft3
Example 1-12: One thousand cubic feet of methane is to be compressed from 60°F and atmospheric pressure to 500 psig and a temperature of 50°F. What volume will it occupy at these conditions?
Solution:
Moles CH4 (n) = 1000 / 379.5 = 2.64
At final conditions, (Compressibility factor z must be calculated), from equations 1-14 and1-15
Tr = (460 + 50) / 344 = 1.88
Pr = (500 + 14.7) / 673 = 0.765
From Figure 1-10, Z = 0.94
From eqn. 1-10, PV = nzRT
V = ft3
Example 1-13: One pound-mole of C3H8 (44 lb) is held in a container having a capacity of 31.2 cu ft. The temperature is 280°F. "What is the pressure?
Solution:
Volume = V = 31.2 ft3
A Trial-and-error solution is necessary because the compressibility factor Z is a function of the unknown pressure. Assume Z = 0.9.
Using Eq. 1-10, PV = nzRT
P ×31.2 = 0.9 × 1.0 × 10.7 × (460 + 280)
P = 229 psia
From table 1-3, eqns. 1-14 and 1-15
Pr = 229 / 616 = 0.37,
Tc = 665ºR
Tr = (460 + 280) / 665 = 1.113
According to Figure 1.9, the value of Z should be about 0.915 rather than 0.9. Thus, recalculate using eq. 1-10, the pressure is 232 rather than 229 psia.
Example 1-14: Calculate the volume of gas (MW=20) will occupy a vessel with diameter 24 in, and 6 ft. length. At pressure 200 psia and temperature 100 0F. (Assume compressibility factor z=0.9), and what will be the volume of gas at 14.7 psia and 60 0F.
Then calculate gas density and mass inside the container at pressure 200 psia and temperature 100 0F.
Volume of vessel = π L r2
V = 3.14 × 6 × (24)2/ (2 ×12)2 ft3
V = 18.8 ft3.
(We divided by 2 to get (r) from the diameter, and divided by 12 to convert from in. to ft.)
T = 460 + 100 = 560 0R
Using Eq. 1-10, PV=nZRT
n = 18.8 × 200 / (0.9 × 10.73 × 560)
n = 0.7 lb. moles. (Remember gas volume ft3 = 379.5 x n)
Volume of gas at 200 psia and 100 0F= 0.7 * 379.5 = 266 ft3
n of Gas at 14.7 psia and 60 0F ( z=1) = 18.8 × 14.7 / (1 × 10.73 × 520)
n = 0.0495 lb. moles
Volume of gas at 14.7 psia and 60 0F = 0.0495 * 379.5 = 18.8 ft3
From the previous example 1-14, the gas volume will equal to the container volume at standard conditions (14.7 psia and 60 0F).
Gas density is calculated using Eq. 1-19
ρg = 0.093 ((MW)P)/TZ lb/ft3
Density of gas ρg = 0.093 × 20 × 200 / (0.9 × 560) = 0.738 lb/ft3
Mass of gas inside the vessel = Volume × density = 0.738 × 265 = 196 lb mass
1.9.4 Velocity of gas, (ft/s)
The velocity of gas equal the volume flow rate (ft3) per second divided by flow area (ft2).
Example 1-15: Calculate the gas velocity for gas flow rate 100 MMscfd through 24 in. internal diameter gas pipe, the gas specific gravity is 0.7, pressure 500 psia, Temperature 100 0F, and assume compressibility factor 0.85.
Solution: Using Eq. 1-10, PV=nzRT, and remember that n= V (ft3)/379.5).
n = 100 × 106/379.5
Gas volume at operating conditions V= 100 × 106 × 0.85 × 10.73 × 560 / (379.5 × 500)
= 2,695,000 ft3/day
Gas flow rate cubic foot per second = 2,695,000 / (24×60×60) = 31.2 ft3/sec
Area of flow = π r2 = 3.14 × 12 × 12 / (144) = 3.14 ft2
(144 to convert r2 from in. to ft2.)
Velocity of gas will be 31.2/3.14 = 9.9 ft/s
The gas velocity may be calculated directly from the following equation:
Velocity = 6 ZTQ/(100,000×Pd2) ft/s. Eq 1-20
Where Q = Flow rate scfd, d = diameter in inches.
The maximum recommended velocity of dry gas in pipes is 100 ft/s, (60 ft/s for wet gas), and to be less than the erosional velocity which is defined as:
Erosional velocity: The erosional velocity represents the upper limit of gas velocity in a pipeline. As the gas velocity increases, vibration and noise result. Higher velocities also cause erosion of the pipe wall over a long time period. The erosional velocity Vmax may be calculated approximately as follows:
Vmax = 100 √(2&ZRT/29GP) Eq 1-21
Where G= gas sp. Gt (air=1), P = pressure psia
For Example 1-15, the erosional velocity Vmax is:
Vmax = 100 √(2&0.85×10.73×560/(29×0.7× 500)) Vmax = 70.9 ft/s.
1.9.5 Average pipeline pressure
The gas compressibility factor Z used in the General Flow equation is based upon the flowing temperature and the average pipe pressure. The average pressure may be approximated as the arithmetic average
Pavg = (P1+P2)/2 of the upstream and downstream pressures P1 and P2. However, a more accurate average pipe pressure is usually calculated as follows
Pavg = 2/3 (P1+P2 - (P1× P2)/(P1+ P2)) Eq 1-22
Where
P1, P2, Pavg = pressure, psia
Example 1-16: A natural gas pipeline with internal diameter 19 in. transports natural gas (Sp. Gr.= 0.65) at a flow rate of 200 MMscfd. Calculate the gas velocity at inlet and outlet of the pipe, assuming isothermal flow. The inlet temperature of 70 0F, inlet pressure is 1200 psig, and outlet pressure is 900 psig. Use compressibility factor of 0.95. Also, calculate the erosional velocity for this pipeline.
Solution:
Using Eq. 1-20, the gas velocity at inlet of the pipe:
Velocity = 6 × 0.95× 530×200,000,000/(100,000×1214.7×192) ft/s.
Velocity = 13.8 ft/s.
The gas velocity at outlet of the pipe:
Velocity = 6 × 0.95× 530×200,000,000/(100,000×914.7×192) ft/s.
Velocity = 18.3 ft/s.
Finally, the erosional velocity can be calculated using Eq. 1-21
Vmax = 100 √(2&0.95 ×10.73×530/29×0.65×1214.7)
Vmax = 48.6 ft/s.
The above example may be solved by calculating the gas density at inlet and outlet of the pipe, then calculating the operational flow rate, divide it by pipe cross sectional area to get the velocity as follows:
Gas molecular weight = 0.65 × 28.96 = 18.8
Using Eq. 1-10, PV = nzRT
Calculating n = 200,000,000 / 379.5
Flow rate under operating conditions =
Gas volume V (= flow rate Q) = 200,000,000 × 0.95 × 10.73 × 530 / (379.5 ×1214.7)
Q = 2,347,000 ft3 per day at operating conditions. Q = 27.16 ft3/s.
Pipe cross sectional area = π r2 = 3.14 × 19 × 19 /(4× 144) = 1.97 ft2
Velocity of gas at the inlet = 27.16/1.97 = 13.8 ft/s.
1.9.6 Viscosity of gases
Viscosity of a fluid relates to the resistance to flow of the fluid. Higher the viscosity, more difficult it is to flow. Viscosity is a number that represents the drag forces caused by the attractive forces in adjacent fluid layers. It might be considered as the internal friction between molecules, separate from that between the fluid and the pipe wall.
The viscosity of a gas is very small compared to that of a liquid. For example, a typical crude oil may have a viscosity of 10 centipoise (cp), whereas a sample of natural gas has a viscosity of 0.0019 cp.
Viscosity may be referred to as absolute or dynamic viscosity measured in cp or kinematic viscosity measured in centistokes (cSt). Other units of viscosity are lb/ft-sec for dynamic viscosity and ft2/s for kinematic viscosity.
Fluid viscosity changes with temperature. Liquid viscosity decreases with increasing temperature, whereas gas viscosity decreases initially with increasing temperature and then increases with further increasing temperature.
Table 1- 16 Viscosity conversion factors
Figure 1-11 can be used to estimate the viscosity of a hydrocarbon gas at various conditions of temperature and pressure if the specific gravity of the gas at standard conditions is known. It is useful when the gas composition is not known. It does not make corrections for H2S, CO2, and N2. It is useful for determining viscosities at high pressure.
Figure 1-11 Hydrocarbon gas viscosity.
Example 1-9: Calculate the pseudo-critical temperature and pressure for the natural gas stream composition given in example 1-2, calculate the compressibility factor, and gas density at 600 psia and 1000F.
Solution:
Table 1-15 solution of Example 1-9.
From the table MW= 21.36
T`c = 451.5 0R
P`c = 667 psia
From Eq. (1-14) and Eq. (1-15)
Tr = T/T`c = (100+460)/451.5 = 1.24
Pr = P/P`c = 600/667 = 0.9
Compressibility factor z could be calculated from figure 1-10, or from Eq. (1-18)
Value from figure, z = 0.83
From Equation 1-15 z = 0.87
For our further calculation we will use the calculated z value [Eq. (1-18)]
Using eq. (1-19) density of gas
ρg = 0.093 ((21.36)600)/(560 ×0.83) = 2.56 lb/ft3
Comparing ρg at standard condition (z=1)
ρg at standard condition = 0.093 (21.36)14.7/(520 ×1) = 0.056 lb/ft3
We can conclude that density increases with pressure while the volume decreases.
1.9.3 Gas volume at any condition of Pressure and temperature
Volume of a gas is the space occupied by the gas. Gases fill the container that houses the gas. The volume of a gas generally varies with temperature and pressure.
Volume of a gas is measured in cubic feet (ft3).
Gas volume are commonly referred to in "standard" or "normal" units.
Standard conditions commonly refers to gas volumes measured at: 60°F and 14.696 psia
The Gas Processors Association (GPA) SI standard molar volume conditions is 379.49 std ft3/lb mol at 60°F, 14.696 psia.
Therefore, each mole (n) contains about 379.5 cubic feet of gas (ft3)at standard conditions.
Therefore, by knowing the values of mass and density at certain pressure and temperature, the volume occupied by gas can be calculated.
Example 1-10: Calculate the volume of a 10 lb mass of gas (Gravity = 0.6) at 500 psig and 80 0F, assuming the compressibility factor as 0.895. The molecular weight of air may be taken as 29 and the base pressure is 14.7 psia.
Solution:
The molecular weight of the gas (MW) = 0.6 x 29 = 17.4
Pressure =500+14.7 = 514.7 psia
Temperature = 80+460 = 540 0R
Compressibility factor z= 0.895
The number of lb moles n is calculated using Eq. (1-2). n=m/(MW)
n = 10/17.4
Therefore n= 0.5747 lb mole
Using the real gas Eq. (1-10), PV=nzRT
(514.7) V = 0.895 x 0.5747 x 10.73 x 540. Therefore, V = 5.79 ft3
Example 1-11: Calculate the volume of 1 lb mole of the natural gas stream given in the previous example at 1200F and 1500 psia (compressibility factor Z = 0.811).
Solution:
Using Eq.(1-10), PV = nzRT
V= 0.811 x 1 x 10.73 x (120+460)/1500. V = 3.37 ft3
Example 1-12: One thousand cubic feet of methane is to be compressed from 60°F and atmospheric pressure to 500 psig and a temperature of 50°F. What volume will it occupy at these conditions?
Solution:
Moles CH4 (n) = 1000 / 379.5 = 2.64
At final conditions, (Compressibility factor z must be calculated), from equations 1-14 and1-15
Tr = (460 + 50) / 344 = 1.88
Pr = (500 + 14.7) / 673 = 0.765
From Figure 1-10, Z = 0.94
From eqn. 1-10, PV = nzRT
V = ft3
Example 1-13: One pound-mole of C3H8 (44 lb) is held in a container having a capacity of 31.2 cu ft. The temperature is 280°F. "What is the pressure?
Solution:
Volume = V = 31.2 ft3
A Trial-and-error solution is necessary because the compressibility factor Z is a function of the unknown pressure. Assume Z = 0.9.
Using Eq. 1-10, PV = nzRT
P ×31.2 = 0.9 × 1.0 × 10.7 × (460 + 280)
P = 229 psia
From table 1-3, eqns. 1-14 and 1-15
Pr = 229 / 616 = 0.37,
Tc = 665ºR
Tr = (460 + 280) / 665 = 1.113
According to Figure 1.9, the value of Z should be about 0.915 rather than 0.9. Thus, recalculate using eq. 1-10, the pressure is 232 rather than 229 psia.
Example 1-14: Calculate the volume of gas (MW=20) will occupy a vessel with diameter 24 in, and 6 ft. length. At pressure 200 psia and temperature 100 0F. (Assume compressibility factor z=0.9), and what will be the volume of gas at 14.7 psia and 60 0F.
Then calculate gas density and mass inside the container at pressure 200 psia and temperature 100 0F.
Volume of vessel = π L r2
V = 3.14 × 6 × (24)2/ (2 ×12)2 ft3
V = 18.8 ft3.
(We divided by 2 to get (r) from the diameter, and divided by 12 to convert from in. to ft.)
T = 460 + 100 = 560 0R
Using Eq. 1-10, PV=nZRT
n = 18.8 × 200 / (0.9 × 10.73 × 560)
n = 0.7 lb. moles. (Remember gas volume ft3 = 379.5 x n)
Volume of gas at 200 psia and 100 0F= 0.7 * 379.5 = 266 ft3
n of Gas at 14.7 psia and 60 0F ( z=1) = 18.8 × 14.7 / (1 × 10.73 × 520)
n = 0.0495 lb. moles
Volume of gas at 14.7 psia and 60 0F = 0.0495 * 379.5 = 18.8 ft3
From the previous example 1-14, the gas volume will equal to the container volume at standard conditions (14.7 psia and 60 0F).
Gas density is calculated using Eq. 1-19
ρg = 0.093 ((MW)P)/TZ lb/ft3
Density of gas ρg = 0.093 × 20 × 200 / (0.9 × 560) = 0.738 lb/ft3
Mass of gas inside the vessel = Volume × density = 0.738 × 265 = 196 lb mass
1.9.4 Velocity of gas, (ft/s)
The velocity of gas equal the volume flow rate (ft3) per second divided by flow area (ft2).
Example 1-15: Calculate the gas velocity for gas flow rate 100 MMscfd through 24 in. internal diameter gas pipe, the gas specific gravity is 0.7, pressure 500 psia, Temperature 100 0F, and assume compressibility factor 0.85.
Solution: Using Eq. 1-10, PV=nzRT, and remember that n= V (ft3)/379.5).
n = 100 × 106/379.5
Gas volume at operating conditions V= 100 × 106 × 0.85 × 10.73 × 560 / (379.5 × 500)
= 2,695,000 ft3/day
Gas flow rate cubic foot per second = 2,695,000 / (24×60×60) = 31.2 ft3/sec
Area of flow = π r2 = 3.14 × 12 × 12 / (144) = 3.14 ft2
(144 to convert r2 from in. to ft2.)
Velocity of gas will be 31.2/3.14 = 9.9 ft/s
The gas velocity may be calculated directly from the following equation:
Velocity = 6 ZTQ/(100,000×Pd2) ft/s. Eq 1-20
Where Q = Flow rate scfd, d = diameter in inches.
The maximum recommended velocity of dry gas in pipes is 100 ft/s, (60 ft/s for wet gas), and to be less than the erosional velocity which is defined as:
Erosional velocity: The erosional velocity represents the upper limit of gas velocity in a pipeline. As the gas velocity increases, vibration and noise result. Higher velocities also cause erosion of the pipe wall over a long time period. The erosional velocity Vmax may be calculated approximately as follows:
Vmax = 100 √(2&ZRT/29GP) Eq 1-21
Where G= gas sp. Gt (air=1), P = pressure psia
For Example 1-15, the erosional velocity Vmax is:
Vmax = 100 √(2&0.85×10.73×560/(29×0.7× 500)) Vmax = 70.9 ft/s.
1.9.5 Average pipeline pressure
The gas compressibility factor Z used in the General Flow equation is based upon the flowing temperature and the average pipe pressure. The average pressure may be approximated as the arithmetic average
Pavg = (P1+P2)/2 of the upstream and downstream pressures P1 and P2. However, a more accurate average pipe pressure is usually calculated as follows
Pavg = 2/3 (P1+P2 - (P1× P2)/(P1+ P2)) Eq 1-22
Where
P1, P2, Pavg = pressure, psia
Example 1-16: A natural gas pipeline with internal diameter 19 in. transports natural gas (Sp. Gr.= 0.65) at a flow rate of 200 MMscfd. Calculate the gas velocity at inlet and outlet of the pipe, assuming isothermal flow. The inlet temperature of 70 0F, inlet pressure is 1200 psig, and outlet pressure is 900 psig. Use compressibility factor of 0.95. Also, calculate the erosional velocity for this pipeline.
Solution:
Using Eq. 1-20, the gas velocity at inlet of the pipe:
Velocity = 6 × 0.95× 530×200,000,000/(100,000×1214.7×192) ft/s.
Velocity = 13.8 ft/s.
The gas velocity at outlet of the pipe:
Velocity = 6 × 0.95× 530×200,000,000/(100,000×914.7×192) ft/s.
Velocity = 18.3 ft/s.
Finally, the erosional velocity can be calculated using Eq. 1-21
Vmax = 100 √(2&0.95 ×10.73×530/29×0.65×1214.7)
Vmax = 48.6 ft/s.
The above example may be solved by calculating the gas density at inlet and outlet of the pipe, then calculating the operational flow rate, divide it by pipe cross sectional area to get the velocity as follows:
Gas molecular weight = 0.65 × 28.96 = 18.8
Using Eq. 1-10, PV = nzRT
Calculating n = 200,000,000 / 379.5
Flow rate under operating conditions =
Gas volume V (= flow rate Q) = 200,000,000 × 0.95 × 10.73 × 530 / (379.5 ×1214.7)
Q = 2,347,000 ft3 per day at operating conditions. Q = 27.16 ft3/s.
Pipe cross sectional area = π r2 = 3.14 × 19 × 19 /(4× 144) = 1.97 ft2
Velocity of gas at the inlet = 27.16/1.97 = 13.8 ft/s.
1.9.6 Viscosity of gases
Viscosity of a fluid relates to the resistance to flow of the fluid. Higher the viscosity, more difficult it is to flow. Viscosity is a number that represents the drag forces caused by the attractive forces in adjacent fluid layers. It might be considered as the internal friction between molecules, separate from that between the fluid and the pipe wall.
The viscosity of a gas is very small compared to that of a liquid. For example, a typical crude oil may have a viscosity of 10 centipoise (cp), whereas a sample of natural gas has a viscosity of 0.0019 cp.
Viscosity may be referred to as absolute or dynamic viscosity measured in cp or kinematic viscosity measured in centistokes (cSt). Other units of viscosity are lb/ft-sec for dynamic viscosity and ft2/s for kinematic viscosity.
Fluid viscosity changes with temperature. Liquid viscosity decreases with increasing temperature, whereas gas viscosity decreases initially with increasing temperature and then increases with further increasing temperature.
Table 1- 16 Viscosity conversion factors
Figure 1-11 can be used to estimate the viscosity of a hydrocarbon gas at various conditions of temperature and pressure if the specific gravity of the gas at standard conditions is known. It is useful when the gas composition is not known. It does not make corrections for H2S, CO2, and N2. It is useful for determining viscosities at high pressure.
Figure 1-11 Hydrocarbon gas viscosity.
Fundamentals of Oil and Gas Processing
Basics of Gas Field Processing
Basics of Corrosion in Oil and Gas Industry
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Basics of Gas Field Processing
Basics of Corrosion in Oil and Gas Industry
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https://www.hurras.org/vb/
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yasserkassem
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Re: Basics of Gas Field Processing Book "Full text"
------------Fundamentals of Oil and Gas Processing Book
Basics of Gas Field Processing Book
Prediction and Inhibition of Gas Hydrates Book
Basics of Corrosion in Oil and Gas Industry Book
https://www.amazon.com/Yasser-Kassem/e/ ... scns_share
Chapter 2 45
Phase Behavior and Phase separation 45
2.1: Phase Behavior 45
2.1.1: Introduction 45
2.1.2: System Components 45
2.1.4: Single-Component Systems 46
2.1.5: Multicomponent Systems 49
2.1.6: Prediction of phase envelope 50
------------
Chapter 2
Phase Behavior and Phase separation
2.1: Phase Behavior
2.1.1: Introduction
Before studying the separation of gases and liquids, we need to understand the relationship between the phases. Phase defines any homogeneous and physically distinct part of a system that is separated from other parts of the system by definite bounding surfaces:
The matter has three phases, the simplest example is water.
• Solid (ice),
• Liquid (liquid water),
• Vapor (water vapor).
Solids have a definite shape and are hard to the touch. They are composed of molecules with very low energy that stay in one place even though they vibrate. Liquids have a definite volume but no definite shape. Liquids assume the shape of the container but will not necessarily fill that container. Liquid molecules possess more energy than a solid (allows movement from place to another). By virtue of the energy, there is more space between molecules, and liquids are less dense than solids. Vapors do not have a definite volume or shape and will fill a container in which they are placed. Vapor molecules possess more energy than liquids (very active) and are less dense than liquids.
Our primary concern in this section is the difference in energy level between phases.
Energy is added to melt a solid to form a liquid. Additional energy will cause the liquid to vaporize. One needs to know the phase or phases that exist at given conditions of pressure, volume, and temperature so as to determine the corresponding energy level, to do this we need to study the phase diagram or phase behavior, but first we have to separate components into three classifications:
• Pure substance (single-component systems),
• Two substances,
• Multicomponent.
Phase diagrams illustrate the phase that a particular substance will take under specified conditions of pressure, temperature, and volume.
2.1.2: System Components
Natural gas systems are composed primarily of the lighter alkane series of hydrocarbons, with methane (CH4) and ethane (C2H6) comprising 80% to 90% of the volume of a typical mixture. Methane and ethane exist as gases at atmospheric conditions.
Propane (C3H8), butane (n-C4H10 and i-C4H10), and heavier hydrocarbons may be extracted from the gas system and liquefied for transportation and storage. These are the primary components of liquefied petroleum gas, or LPG.
The intermediate-weight hydrocarbons (pentane through decane) exist as volatile liquids at atmospheric conditions. These components are commonly referred to as pentanes-plus, condensate, natural gasoline, and natural gas liquids (NGL).
Natural gas systems can also contain non-hydrocarbon constituents, including hydrogen sulfide (H2S), carbon dioxide (CO2), nitrogen (N2), and water vapor. These constituents may occur naturally in gas reservoirs, or they may enter the system as contaminants during production, processing, and transportation. In addition, operators may intentionally add odorants, tracers (such as helium), or other components.
Dry, or lean, natural gas systems have high concentrations of the lighter hydrocarbons (methane and ethane), while wet, or rich, gas systems have higher concentrations of the intermediate-weight hydrocarbons. Lean gases burn with a low air-to-gas ratio and display a colorless to blue or yellow flame, whereas rich gases require comparatively higher amounts of air for combustion and burn with an orange flame. Intermediate-weight hydrocarbons may condense from rich gases upon cooling.
2.1.4: Single-Component Systems
A pure component of a natural gas system exhibits a characteristic phase behavior, as shown in Fig. 2-1. Depending on the component’s pressure and temperature, it may exist as a vapor, a liquid, or some equilibrium combination of vapor and liquid
Figure 2-1 P-T Diagram for pure component.
Lines HD, HC, and FH are the equilibrium lines - combinations of pressure and temperature at which the adjoining phases are in equilibrium. At equilibrium, one can change phase, by simply adding or removing energy from the system. Point H, the triple point, is the only combination of pressure and temperature at which all three phases can exist together.
Along line FH no liquid phase is ever present and solid sublimes to vapor. The use of "dry ice" for cooling is an example of this. Line HD is the equilibrium line between solid and liquid. Ice water at 0°C [32°F] and atmospheric pressure occurs on this line. Line HD can have a positive or negative slope depending on whether the liquid expands or contracts on freezing. The energy change occurring along line HD is called the heat of fusion. At any P and T along this line the system can be all solid, all liquid or a mixture of the two depending on the energy level.
This line could be called the solid-liquid saturation or solid-liquid equilibrium line.
Line HC is the saturation or equilibrium curve between vapor and liquid. It starts at the triple point and terminates at the critical point "C." The pressure and temperature conditions at this latter point are known as critical temperature (Tc) and critical pressure (Pc).
At this point the properties of the liquid and vapor phases become identical. For a pure substance the critical point can be defined as that point above which liquid cannot exist as a unique separate phase. Above (Pc), and (Tc), the system is often times referred to as a dense fluid to distinguish it from normal vapor and liquid.
Line HC is often referred to as the vapor pressure curve. Such vapor pressure curves are available from many sources. Line HC is also the bubble point and dew point curve for the pure substance.
The vapor pressure line in Figure 2-2 divides the liquid region from the vapor region.
In figure 2-1, consider a process starting at pressure P1, and proceeding at constant pressure.
From "m" to "n" the system is entirely solid. The system is all liquid for the segment o-b. At "b" the system is a saturated liquid - any further addition of energy will cause vaporization. At "d," the system is in the saturated vapor state. At temperatures above "d," it is a superheated vapor.
Line HC is thus known by many names - equilibrium, saturated, bubble point, dew point and vapor pressure. For a pure substance these words all mean the same thing.
At the pressure and temperature represented by HC the system may be all saturated liquid, all saturated vapor or a mixture of vapor and liquid.
The rectangle "bfghd" illustrates another important phase property that is confirmed experimentally.
Suppose we place a liquid in a windowed cell at condition "b" and light it so it is easily visible.
We then increase pressure at constant temperature (isothermally). As we proceed toward point “f” the color will begin to fade. At some point, the color disappears completely. The cell now contains what looks like a vapor, but no bubble of vapor was ever seen to form.
At “ f ” (above the critical) the system is in a fourth phase that cannot be described by the senses. It is usually called dense phase fluid, or simply fluid. The word "fluid" refers to anything that will flow and applies equally well to gas and liquid.
This fluid at "f' looks like a gas but possesses different properties from regular gas found to the right of line HC and below the critical pressure. It is denser than regular gas but is more compressible than a regular liquid. “Properties of the liquid and vapor phases become identical”.
Table 1-3 lists Critical pressures and critical temperatures, along with molecular weights, of some pure components present in many natural gas systems.
Figure 2- 2 shows vapor pressure line for light hydrocarbons, where the left part of any component line, represents its liquid phase while the right part represents its gas phase.
Figure 2-2 Vapor pressure for light hydrocarbons.
2.1.5: Multicomponent Systems
In reality, natural gas systems are not pure substances. Rather, they are mixtures of various components, with phase behavior characteristics that differ from those of a single-component system. Instead of having a vapor pressure curve, a mixture exhibits a phase envelope, as shown in Figure 2-3.
Figure 2-3 typical phase envelop of hydrocarbon mixture.
The phase envelope (curve BCD in Figure 2-3) separates the liquid and gas phases. The area within this envelope is called the two-phase region and represents the pressure and temperature ranges at which liquid and gas exist in equilibrium.
The upper line of the two-phase region (curve BC) is the bubble-point line. This line indicates where the first bubble of vapor appears when the pressure of the liquid phase mixture is lowered at constant temperature, or when the temperature increases at constant pressure.
The lower section of the phase envelope (curve CD) is the dewpoint line. When the pressure of a mixture in the gaseous phase is decreased at constant temperature, or when the temperature is lowered at constant pressure, the first drop of liquid forms on this line. The bubble-point line and the dewpoint line meet at the critical point (C).
The highest pressure in the two-phase region is called the cricondenbar, while the highest temperature in the two-phase region is called the cricondentherm.
Figure 2-4, is another example of phase envelope, where:
Cricondenbar - maximum pressure at which liquid and vapor may exist (Point N).
Cricondentherm - maximum temperature at which liquid and vapor may coexist in equilibrium (Point M).
Retrograde Region - that area inside phase envelope where condensation of liquid occurs by lowering pressure or increasing temperature (opposite of normal behavior).
Quality Lines - those lines showing constant percentages which intersect at the critical point (C) and are essentially parallel to the bubble point and dew point curves. The bubble point curve represents 0% vapor and the dew point curve 100% vapor.
Line ABDE represents a typical isothermal retrograde condensation process occurring in a condensate reservoir. Point A represents the single phase fluid outside the phase envelope. As pressure is lowered, Point B is reached where condensation begins. As pressure is lowered further, more liquid forms because of the change in the slope of the quality lines. As the process continues outside the retrograde area, less and less liquid forms until the dewpoint is reached (Point E). Below E no liquid forms.
Figure 2-4 shows another phase envelope for hydrocarbon mixture.
2.1.6: Prediction of phase envelope
The location of the bubblepoint and dewpoint lines may be calculated using vapor-liquid equilibrium (VLE) methods. For most naturally occurring systems above about [2000 psia], the validity of the standard calculation becomes questionable.
The application of K-values to calculate phase quantities and compositions proceeds as follows.
For any stream (F) with mole fractions of components (Z1+Z2+Z3,.., etc.) entering a vessel at certain pressure and temperature, the stream will be divided into Vapor stream(V) with mole fractions of components (Y1+Y2+Y3,.., etc.), and into a liquid phase (L) with mole fractions of components (X1+X2+X3,.., etc.).
Component balance:
Fzi = Vyi + Lxi Eq. 2-1
Figure 2-5 flash separation for hydrocarbon mixture.
where
zi = mol fraction of any component in total feed stream to separation vessel
yi = mol fraction of any component in the vapor phase
xi = mol fraction of any component in the liquid phase
Ki = equilibrium vaporization ratio (equilibrium constant) = yi/xi
F = total mols of feed
V = total mols of vapor
L = total mols of liquid
If we set F = 1.0 so that L and V are now liquid and vapor-to-feed ratios
then zi = Vyi + Lxi
Since yi = Kixi
So, zi = V Ki xi + L xi
xi = zi / ( L + V Ki) Eq. 2-2
Since the summation of liquid fractions must equal one, we can write the following equation.
∑ xi = ∑ zi / ( L + V Ki) = 1 Eq. 2-3
The equation serves as the objective function in an interactive calculation to determine the quantity of L or V. The calculation procedure is as follows:
1. Determine K values of each component at the temperature and pressure of the system.
2. Assume a value of L (remember, V = 1 - L)
3. Solve the equation Eq. 2-3. If ∑xi ≠ 1.0 assume a new value of L and repeat step 2.
4. When ∑ xi = 1.00, the phase quantities L and V are known as well as the liquid phase composition. Vapor phase compositions may be calculated by remembering that yi = Kixi,
The foregoing calculations is known as a flash calculation and is used to predict the equilibrium quantities and compositions of two phase systems.
Special cases of a flash calculation include bubble point (V = 0, L = 1) and dew point (V = 1, L = 0), calculations. Equations for bubblepoint, and dewpoint are as follows:
Bubblepoint condition:
∑ Ki xi = 1.0 Eq. 2-4
Dewpoint condition:
∑ yi/Ki = 1.0 Eq. 2-5
Flash calculation are usually made by computer software, but knowing the basic of calculations is important in understanding the gas-liquid separation process.
Example 2- 1: Calculate the bubblepoint and dewpoint temperature at 250 psia of the following hydrocarbon mixture. Then calculate the amount of vapor and liquid and the composition of the two phases if these feed entered a vessel @ 250 psia and 150 0F.
Table 2-1 hydrocarbon component for example 2-1.
Solution:
Bubblepoint calculation : To calculate the bubblepoint temperature at certain pressure, (All the components are in liquid phase xi = 1).
From eq. 2-4, the bubblepoint will be reached when ∑ Ki xi ≅ 1
Solution Steps:
Assume a temperature value (100 0F), for example.
From the K chart of each compound, find the K value at the system pressure and assumed temperature.
Multiply mole fraction xi of each component by its equilibrium value taken from the table Ki.
Take the sum ∑ Ki xi , if it’s less than 1, choose higher temperature, (150 0F for example), and repeat as in the table.
If, ∑ Ki xi is higher than 1, choose a lower temperature.
Repeat till ∑ Ki xi ≅ 1.
Table 2-2 bubblepoint calculation for example 2-1.
We assumed two values of temperature , we found the first value (100 0F) is lower than the bubble point since ∑ Ki xi < 1.00 , and the second value (150 0F) is higher than the bubble point since ∑ Ki xi > 1.00 , the bubble point will be between the two values where ∑ Ki xi ≅ 1.
The Ki values in previous table where collected from “ Design operation and maintenance of gas plants - John Campbell Co.” , since it’s hard to obtain Ki numbers at temperature rather than the pre-drawn temperature lines in K-Charts.
The Values of Ki can be extracted from individual component charts (figures 2-6 to 2-10) (Methane K-chart, Ethane K-chart ….etc.), or can be extracted from “DePriester” chart, fig 2-11.
Dewpoint calculation: To calculate the dewpoint temperature at certain pressure, (All components are in gas phase yi = 1). From eq. 2-5, the dewpoint will be reached when ∑ yi /Ki ≅ 1
Solution Steps:
Assume a temperature value (150 0F), for example.
From the K chart of each compound, find the K value at the system pressure and assumed temperature.
Divide mole fraction Yi of each component by its equilibrium value taken from the table Ki.
Take the sum ∑ Yi /Ki , if it’s higher than 1, choose higher temperature, (2000F for example), and repeat as in the table.
If, ∑ Yi /Ki is less than 1, choose a lower temperature.
Repeat till ∑ Yi /Ki ≅ 1.
Table 2-3 dewpoint calculation for example 2-1.
We assumed two values of temperature , we found the first value (150 0F) is lower than the dewpoint since ∑ yi /Ki > 1.00 , and the second value (200 0F) is higher than the bubble point since ∑ yi /Ki value is < 1.00 , the dewpoint will be between the two values where ∑ yi /Ki ≅ 1 .
The Ki values in previous table where collected from “ Design operation and maintenance of gas plants - John Campbell Co.” , since it’s hard to obtain Ki numbers at temperature rather than the pre-drawn temperature lines in K-Charts.
The Values of Ki can be extracted from individual component charts (Methane K-chart, Ethane K-chart ….etc.), or can be extracted from “DePriester” chart, fig 2-11.
Flash calculations:
Different values of “L” will be assumed (remember, V = 1 - L), and accordingly Xi will be calculated till we obtain ∑ xi = 1.00.
Ki from chart at 250 psia and 150 0F
Using Eq. 2 -2 xi = zi / ( L + V Ki)
Component Zi Ki Assume
L=0.5
Xi = Assume
L = 0.75
Xi = Assume
L= 0.649
Xi =
Yi = Ki Xi
Table 2-4 flash calculations for example 2-1.
The assumed value of L=0.5, found to be lower than the correct value, and the assumed value of L= 0.75 found to be higher than the correct value.
The correct value must be in between the two previous assumed values, and found to be 0.649.
Flash calculations usually performed by computer software, for manual calculations, some K value charts are included in this chapter for the illustration of manual calculations for the previous example. (Figures 2-6 to 2-10)
Other K-values are included in Chapter 25 “Equilibrium Ratio (K) Data” in the “GPSA Engineering Data Book”, or Appendix 5A Volume 1 “Gas conditioning and Processing – The Basic Principles,” Campbell Petroleum Series. In the other hand, the DePriester Chart Figure 2-11, may be used for all hydrocarbon components.
K Value charts:
Figure 2-6 Equilibrium ratio (K) for Methane.
Figure 2-7 Equilibrium ratio (K) for Ethane.
Figure 2-8 Equilibrium ratio (K) for Propane.
Figure 2-9 Equilibrium ratio (K) for i-Butane.
Figure 2-10 Equilibrium ratio (K) for n-Butane.
Figure 2-11 the DePriester (K) Chart for hydrocarbon components.
Fundamentals of Oil and Gas Processing
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Re: Basics of Gas Field Processing Book "Full text"
Chapter 3
Detailed Two-phase Oil and Gas Separation can be found in Fundamentals of oil and gas processing
http://www.fanarco.net/bb3/viewtopic.php?f=8&t=6176
Chapter 2 .
----------------
Chapter 3 61
Two-phase Oil and Gas Separation 61
3.1 Introduction 61
3.2 Phase Equilibrium 61
3.3: Separation process: 62
3.4: Principles of Physical Separation: 62
3.5: Gravity Separation: 62
3.6: Stage Separation 65
3.6.1: Initial Separation Pressure 65
3.6.2: Stage Separation 66
3.6.3: Selection of Stages 67
3.6.4: Fields with Different Flowing Tubing Pressures 68
3.6.5: Determining Separator Operating Pressures 68
---------------
Chapter 3
Two-phase Oil and Gas Separation
3.1 Introduction
The production system begins at the wellhead. Fluids produced from oil and gas wells generally constitute mixtures of crude oil, natural gas, and salt water. Crude oil–gas–water mixtures produced from wells, are generally directed, through flow lines and manifold system, to a central processing and treatment facility normally called the gas–oil separation plant (GOSP).
The goal is to attain in the downstream (output) of the “gas oil separation plant”, the following components:
Oil free of water and meets other purchaser’s specifications.
Gas free of hydrocarbon liquid meets other purchaser’s specifications.
Water free of oil and meets environmental, and reservoir regulation for disposal or reinjection.
The first step in processing of the produced stream is the separation of the phases (oil, gas, and water) into separate streams.
Oil may still contain between 10% and 15% water that exists mostly as emulsified water, once initial separation is done, each stream undergoes the proper processing for further field treatment.
3.2 Phase Equilibrium
Equilibrium is a theoretical condition that describes an operating system that has reached a “steady-state” condition whereby the vapor is condensing to a liquid at exactly the same rate at which liquid is boiling to vapor. Simply stated, phase equilibrium is a condition where the liquids and vapors have reached certain pressure and temperature conditions at which they can separate. In most production systems, true equilibrium is never actually reached; however, vapors and liquids move through the system slow enough that a “pseudo” or “quasi” equilibrium is assumed. This assumption simplifies process calculations.
Figure 2-1 illustrates several operating points on a generic phase equilibrium diagram. Point A represents the operating pressure and temperature in the petroleum reservoir. Point B represents the flowing conditions at the bottom of the production tubing of a well. Point C represents the flowing conditions at the wellhead. Typically, these conditions are called flowing tubing pressure (FTP) and flowing tubing temperature (FTT). Point D represents the surface conditions at the inlet of the first separator.
Figure 3-1 Phase equilibrium phase diagram for a typical production system.
3.3: Separation process:
The process can be described as:
Two phase separation, or
Three phase separation
The phases referred to are oil, water and gas.
In two phase separation, gas is removed from total liquid (oil plus water).
In three phase separation, however, in addition to the removal of gas from liquids, the oil and water are separated from each other.
Figure 3.2 shows the difference between 2 and 3 phase separation.
3.4: Principles of Physical Separation:
Three principles used to achieve physical separation of gas and liquids or solids are momentum, gravity settling, and coalescing.
Any separator may employ one or more of these principles, but the fluid phases must be "immiscible" and have different densities for separation to occur.
3.5: Gravity Separation:
Since a separation depends upon gravity to separate the fluids, the ease with which two fluids can be separated depends upon the difference in the density or weight per unit volume of the fluids. (Density of liquid is much higher than density of gases).
In the process of separating, separation stages are as follows:
1- Separate liquid mist from the gas phase.
2- Separate gas in the form of foam from the liquid phase.
3- In case of 3 phase separation, in addition to the above two requirements, water droplets should be separated from oil phase, and oil droplets should be separated from water phase.
Figure 3.2 The Difference between 2 & 3 Phase Separation.
Droplets of liquid mist will settle out from gas, provided:
The gas remains in the separator long enough for mist to drop out.
The flow of the gas through the separator is slow enough that no turbulence occurs, which will keep the gas stream stirred up so that the liquid has no chance to drop out.
The objective of ideal two-phase separation, is to separate the hydrocarbon stream into liquid-free gas and gas-free-liquid. Ideally, the gas and liquid reach a state of equilibrium at the existing conditions of Pressure and Temperature within the vessel.
Liquid droplets will settle out of a gas phase due to the difference in densities if the gravitational force acting on the droplet is greater than the drag force of the gas flowing around the droplet (see Fig. 2-3). The drag force is the force resulted from the velocity of gas and affecting the entrained droplet of liquid, forcing it to move in the gas flow direction.
Fig. 3-3 A schematic of a force balance on a droplet in a flowing gas stream.
Figures 3-4, and 3-5, illustrates the liquid droplet in gas phase and gas bubble in liquid phase in both configurations of horizontal and vertical separators.
From both figures, it’s clear that, in vertical separator, the gravitational settling force is countercurrent or opposite of the drag force resulted from gas movement. While in horizontal separator, the two forces are perpendicular to each other.
The same for the gas bubble entrained in liquid in vertical and horizontal separators.
Fig. 3- 4.The liquid droplet in gas phase and gas bubble in liquid phase in horizontal separator.
Fig. 3-5 The liquid droplet in gas phase and gas bubble in liquid phase in vertical separator.
3.6: Stage Separation
3.6.1: Initial Separation Pressure
Because of the multicomponent nature of the produced fluid, the higher the pressure at which the initial separation occurs, the more liquid will be obtained in the separator. This liquid contains some light components that vaporize in the stock tank downstream of the separator. If the pressure for initial separation is too high, too many light components will stay in the liquid phase at the separator and be lost to the gas phase at the tank. If the pressure is too low, not as many of these light components will be stabilized into the liquid at the separator and they will be lost to the gas phase.
This phenomenon, which can be calculated using flash equilibrium techniques discussed in previous chapter, is shown in Figures 3-6 and 3-7.
Fig. 3-6. Single stage separation.
It is important to understand this phenomenon qualitatively. The tendency of any one component in the process stream to flash to the vapor phase depends on its partial pressure. The partial pressure of a component in a vessel is defined as the number of molecules of that component in the vapor space divided by the total number of molecules of all components in the vapor space times the pressure in the vessel [refer to Eq. (3-1)]:
PPN =P × MolesN / ∑ MolesN Eq. 3-1
where
PPN = partial pressure of component “N,”
MolesN = number of moles of component “N,”
Ʃ MolesN = total number of moles of all components,
P = pressure in the vessel, psia.
Thus, if the pressure in the vessel is high, the partial pressure for the component will be relatively high and the molecules of that component will tend toward the liquid phase. This is seen by the top line in Figure 3-7.
As the separator pressure is increased, the liquid flow rate out of the separator increases.
The problem with this is that many of these molecules are the lighter hydrocarbons (methane, ethane, and propane), which have a strong tendency to flash to the gas state at stock-tank conditions (atmospheric pressure). In the stock tank, the presence of these large numbers of molecules creates a low partial pressure for the intermediate-range hydrocarbons (butanes, pentane, and heptane) whose flashing tendency at stock tank conditions is very susceptible to small changes in partial pressure. Thus, by keeping the lighter molecules in the feed to the stock tank, we manage to capture a small amount of them as liquids, but we lose to the gas phase many more of the intermediate-range molecules. That is why beyond some optimum point there is actually a decrease in stock-tank liquids by increasing the separator operating pressure.
Fig. 3-7. Effect of separator pressure on liquid recovery.
3.6.2: Stage Separation
Figure 3-6 deals with a simple single-stage process. That is, the fluids are flashed in an initial separator and then the liquids from that separator are flashed again at the stock tank. Traditionally, the stock tank is not normally considered a separate stage of separation, though it most assuredly is.
Figure 3-8 shows a three-stage separation process. The liquid is first flashed at an initial pressure and then flashed at successively lower pressures two times before entering the stock tank.
Because of the multicomponent nature of the produced fluid, it can be shown by flash calculations that the more stages of separation after the initial separation, the more light components will be stabilized into the liquid phase. This can be understood qualitatively by realizing that in a stage separation process the light hydrocarbon molecules that flash are removed at relatively high pressure, keeping the partial pressure of the intermediate hydrocarbons lower at each stage. As the number of stages approaches infinity, the lighter molecules are removed as soon as they are formed and the partial pressure of the intermediate components is maximized at each stage. The compressor horsepower required is also reduced by stage separation as some of the gas is captured at a higher pressure than would otherwise have occurred. This is demonstrated by the example in Table 3-1.
Table. 3-1. Effect of separation pressure for a rich condensate stream.
Fig. 3-8. Stage separation
3.6.3: Selection of Stages
As shown in Figure 3-9, as more stages are added to the process there is less and less incremental liquid recovery. The diminishing income for adding a stage must more than offset the cost of the additional separator, piping, controls, space, and compressor complexities. It is clear that for each facility there is an optimum number of stages. In most cases, the optimum number of stages is very difficult to determine as it may be different from well to well and it may change as the well’s flowing pressure declines with time. Table 2-7 is an approximate guide to the number of stages in separation, excluding the stock tank, which field experience indicates is somewhat near optimum. Table 3-2 is meant as a guide and should not replace flash calculations, engineering studies, and engineering judgment.
Fig.3-9. Incremental liquid recovery versus number of separator stages.
Table. 3-2. Stage separation guidelines.
3.6.4: Fields with Different Flowing Tubing Pressures
The discussion to this point has focused on a situation where all the wells in a field produce at roughly the same flowing tubing pressure, and stage separation is used to maximize liquid production and minimize compressor horsepower. Often, stage separation is used because different wells producing to the facility have different flowing tubing pressures. This could be because they are completed in different reservoirs, or are located in the same reservoir but have different water production rates. By using a manifold arrangement and different primary separator operating pressures, there is not only the benefit of stage separation of high-pressure liquids, but also conservation of reservoir energy. High-pressure wells can continue to flow at sales pressure requiring no compression, while those with lower tubing pressures can flow into whichever system minimizes compression.
3.6.5: Determining Separator Operating Pressures
The choice of separator operating pressures in a multistage system is large. For large facilities many options should be investigated before a final choice is made. For facilities handling less than 50,000 bpd, there are practical constraints that help limit the options.
A minimum pressure for the lowest-pressure stage would be in the 25- to 50-psig range. This pressure will probably be needed to allow the oil to be dumped to a treater or tank and the water to be dumped to the water treating system. The higher the operating pressure, the smaller the compressor needed to compress the flash gas to sales. Compressor horsepower requirements are a function of the absolute discharge pressure divided by the absolute suction pressure.
Increasing the low-pressure separator pressure from 50 psig to 200 psig may decrease the compression horsepower required by 33%. However, it may also add backpressure to wells, restricting their flow, and allow more gas to be vented to atmosphere at the tank. Usually, an operating pressure of between 50 and 100 psig is optimum.
As stated before, the operating pressure of the highest-pressure separator will be no higher than the sales gas pressure. A possible exception to this could occur where the gas lift pressure is higher than the sales gas pressure. In choosing the operating pressures of the intermediate stages, it is useful to remember that the gas from these stages must be compressed.
Normally, this will be done in a multistage compressor. For practical reasons, the choice of separator operating pressures should match closely and be slightly greater than the compressor inter-stage pressures.
Fig. 3-10. Compressor stages and inlet points of separated gas from multistage separation.
The most efficient compressor sizing will be with a constant compressor ratio per stage. Therefore, an approximation of the intermediate separator operating pressures can be derived from
R = (Pd/Ps)1/n Eq. 3-2
where
R = Compression ratio per stage,
Pd = discharge pressure, psia,
Ps = suction pressure, psia,
n = number of stages.
In order to minimize inter-stage temperatures, the maximum ratio per stage will normally be in the range of 3.6 to 4.0. That means that most production facilities will have either two- or three-stage compressors. A two-stage compressor only allows for one possible intermediate separator operating pressure. A three-stage allows for either one operating at second- or third-stage suction pressure or two intermediate separators each operating at one of the two compressor intermediate suction pressures.( fig. 3-10).
Detailed Two-phase Oil and Gas Separation can be found in Fundamentals of oil and gas processing
http://www.fanarco.net/bb3/viewtopic.php?f=8&t=6176
Chapter 2 .
----------------
------------Fundamentals of Oil and Gas Processing Book
Basics of Gas Field Processing Book
Prediction and Inhibition of Gas Hydrates Book
Basics of Corrosion in Oil and Gas Industry Book
https://www.amazon.com/Yasser-Kassem/e/ ... scns_share
Chapter 3 61
Two-phase Oil and Gas Separation 61
3.1 Introduction 61
3.2 Phase Equilibrium 61
3.3: Separation process: 62
3.4: Principles of Physical Separation: 62
3.5: Gravity Separation: 62
3.6: Stage Separation 65
3.6.1: Initial Separation Pressure 65
3.6.2: Stage Separation 66
3.6.3: Selection of Stages 67
3.6.4: Fields with Different Flowing Tubing Pressures 68
3.6.5: Determining Separator Operating Pressures 68
---------------
Chapter 3
Two-phase Oil and Gas Separation
3.1 Introduction
The production system begins at the wellhead. Fluids produced from oil and gas wells generally constitute mixtures of crude oil, natural gas, and salt water. Crude oil–gas–water mixtures produced from wells, are generally directed, through flow lines and manifold system, to a central processing and treatment facility normally called the gas–oil separation plant (GOSP).
The goal is to attain in the downstream (output) of the “gas oil separation plant”, the following components:
Oil free of water and meets other purchaser’s specifications.
Gas free of hydrocarbon liquid meets other purchaser’s specifications.
Water free of oil and meets environmental, and reservoir regulation for disposal or reinjection.
The first step in processing of the produced stream is the separation of the phases (oil, gas, and water) into separate streams.
Oil may still contain between 10% and 15% water that exists mostly as emulsified water, once initial separation is done, each stream undergoes the proper processing for further field treatment.
3.2 Phase Equilibrium
Equilibrium is a theoretical condition that describes an operating system that has reached a “steady-state” condition whereby the vapor is condensing to a liquid at exactly the same rate at which liquid is boiling to vapor. Simply stated, phase equilibrium is a condition where the liquids and vapors have reached certain pressure and temperature conditions at which they can separate. In most production systems, true equilibrium is never actually reached; however, vapors and liquids move through the system slow enough that a “pseudo” or “quasi” equilibrium is assumed. This assumption simplifies process calculations.
Figure 2-1 illustrates several operating points on a generic phase equilibrium diagram. Point A represents the operating pressure and temperature in the petroleum reservoir. Point B represents the flowing conditions at the bottom of the production tubing of a well. Point C represents the flowing conditions at the wellhead. Typically, these conditions are called flowing tubing pressure (FTP) and flowing tubing temperature (FTT). Point D represents the surface conditions at the inlet of the first separator.
Figure 3-1 Phase equilibrium phase diagram for a typical production system.
3.3: Separation process:
The process can be described as:
Two phase separation, or
Three phase separation
The phases referred to are oil, water and gas.
In two phase separation, gas is removed from total liquid (oil plus water).
In three phase separation, however, in addition to the removal of gas from liquids, the oil and water are separated from each other.
Figure 3.2 shows the difference between 2 and 3 phase separation.
3.4: Principles of Physical Separation:
Three principles used to achieve physical separation of gas and liquids or solids are momentum, gravity settling, and coalescing.
Any separator may employ one or more of these principles, but the fluid phases must be "immiscible" and have different densities for separation to occur.
3.5: Gravity Separation:
Since a separation depends upon gravity to separate the fluids, the ease with which two fluids can be separated depends upon the difference in the density or weight per unit volume of the fluids. (Density of liquid is much higher than density of gases).
In the process of separating, separation stages are as follows:
1- Separate liquid mist from the gas phase.
2- Separate gas in the form of foam from the liquid phase.
3- In case of 3 phase separation, in addition to the above two requirements, water droplets should be separated from oil phase, and oil droplets should be separated from water phase.
Figure 3.2 The Difference between 2 & 3 Phase Separation.
Droplets of liquid mist will settle out from gas, provided:
The gas remains in the separator long enough for mist to drop out.
The flow of the gas through the separator is slow enough that no turbulence occurs, which will keep the gas stream stirred up so that the liquid has no chance to drop out.
The objective of ideal two-phase separation, is to separate the hydrocarbon stream into liquid-free gas and gas-free-liquid. Ideally, the gas and liquid reach a state of equilibrium at the existing conditions of Pressure and Temperature within the vessel.
Liquid droplets will settle out of a gas phase due to the difference in densities if the gravitational force acting on the droplet is greater than the drag force of the gas flowing around the droplet (see Fig. 2-3). The drag force is the force resulted from the velocity of gas and affecting the entrained droplet of liquid, forcing it to move in the gas flow direction.
Fig. 3-3 A schematic of a force balance on a droplet in a flowing gas stream.
Figures 3-4, and 3-5, illustrates the liquid droplet in gas phase and gas bubble in liquid phase in both configurations of horizontal and vertical separators.
From both figures, it’s clear that, in vertical separator, the gravitational settling force is countercurrent or opposite of the drag force resulted from gas movement. While in horizontal separator, the two forces are perpendicular to each other.
The same for the gas bubble entrained in liquid in vertical and horizontal separators.
Fig. 3- 4.The liquid droplet in gas phase and gas bubble in liquid phase in horizontal separator.
Fig. 3-5 The liquid droplet in gas phase and gas bubble in liquid phase in vertical separator.
3.6: Stage Separation
3.6.1: Initial Separation Pressure
Because of the multicomponent nature of the produced fluid, the higher the pressure at which the initial separation occurs, the more liquid will be obtained in the separator. This liquid contains some light components that vaporize in the stock tank downstream of the separator. If the pressure for initial separation is too high, too many light components will stay in the liquid phase at the separator and be lost to the gas phase at the tank. If the pressure is too low, not as many of these light components will be stabilized into the liquid at the separator and they will be lost to the gas phase.
This phenomenon, which can be calculated using flash equilibrium techniques discussed in previous chapter, is shown in Figures 3-6 and 3-7.
Fig. 3-6. Single stage separation.
It is important to understand this phenomenon qualitatively. The tendency of any one component in the process stream to flash to the vapor phase depends on its partial pressure. The partial pressure of a component in a vessel is defined as the number of molecules of that component in the vapor space divided by the total number of molecules of all components in the vapor space times the pressure in the vessel [refer to Eq. (3-1)]:
PPN =P × MolesN / ∑ MolesN Eq. 3-1
where
PPN = partial pressure of component “N,”
MolesN = number of moles of component “N,”
Ʃ MolesN = total number of moles of all components,
P = pressure in the vessel, psia.
Thus, if the pressure in the vessel is high, the partial pressure for the component will be relatively high and the molecules of that component will tend toward the liquid phase. This is seen by the top line in Figure 3-7.
As the separator pressure is increased, the liquid flow rate out of the separator increases.
The problem with this is that many of these molecules are the lighter hydrocarbons (methane, ethane, and propane), which have a strong tendency to flash to the gas state at stock-tank conditions (atmospheric pressure). In the stock tank, the presence of these large numbers of molecules creates a low partial pressure for the intermediate-range hydrocarbons (butanes, pentane, and heptane) whose flashing tendency at stock tank conditions is very susceptible to small changes in partial pressure. Thus, by keeping the lighter molecules in the feed to the stock tank, we manage to capture a small amount of them as liquids, but we lose to the gas phase many more of the intermediate-range molecules. That is why beyond some optimum point there is actually a decrease in stock-tank liquids by increasing the separator operating pressure.
Fig. 3-7. Effect of separator pressure on liquid recovery.
3.6.2: Stage Separation
Figure 3-6 deals with a simple single-stage process. That is, the fluids are flashed in an initial separator and then the liquids from that separator are flashed again at the stock tank. Traditionally, the stock tank is not normally considered a separate stage of separation, though it most assuredly is.
Figure 3-8 shows a three-stage separation process. The liquid is first flashed at an initial pressure and then flashed at successively lower pressures two times before entering the stock tank.
Because of the multicomponent nature of the produced fluid, it can be shown by flash calculations that the more stages of separation after the initial separation, the more light components will be stabilized into the liquid phase. This can be understood qualitatively by realizing that in a stage separation process the light hydrocarbon molecules that flash are removed at relatively high pressure, keeping the partial pressure of the intermediate hydrocarbons lower at each stage. As the number of stages approaches infinity, the lighter molecules are removed as soon as they are formed and the partial pressure of the intermediate components is maximized at each stage. The compressor horsepower required is also reduced by stage separation as some of the gas is captured at a higher pressure than would otherwise have occurred. This is demonstrated by the example in Table 3-1.
Table. 3-1. Effect of separation pressure for a rich condensate stream.
Fig. 3-8. Stage separation
3.6.3: Selection of Stages
As shown in Figure 3-9, as more stages are added to the process there is less and less incremental liquid recovery. The diminishing income for adding a stage must more than offset the cost of the additional separator, piping, controls, space, and compressor complexities. It is clear that for each facility there is an optimum number of stages. In most cases, the optimum number of stages is very difficult to determine as it may be different from well to well and it may change as the well’s flowing pressure declines with time. Table 2-7 is an approximate guide to the number of stages in separation, excluding the stock tank, which field experience indicates is somewhat near optimum. Table 3-2 is meant as a guide and should not replace flash calculations, engineering studies, and engineering judgment.
Fig.3-9. Incremental liquid recovery versus number of separator stages.
Table. 3-2. Stage separation guidelines.
3.6.4: Fields with Different Flowing Tubing Pressures
The discussion to this point has focused on a situation where all the wells in a field produce at roughly the same flowing tubing pressure, and stage separation is used to maximize liquid production and minimize compressor horsepower. Often, stage separation is used because different wells producing to the facility have different flowing tubing pressures. This could be because they are completed in different reservoirs, or are located in the same reservoir but have different water production rates. By using a manifold arrangement and different primary separator operating pressures, there is not only the benefit of stage separation of high-pressure liquids, but also conservation of reservoir energy. High-pressure wells can continue to flow at sales pressure requiring no compression, while those with lower tubing pressures can flow into whichever system minimizes compression.
3.6.5: Determining Separator Operating Pressures
The choice of separator operating pressures in a multistage system is large. For large facilities many options should be investigated before a final choice is made. For facilities handling less than 50,000 bpd, there are practical constraints that help limit the options.
A minimum pressure for the lowest-pressure stage would be in the 25- to 50-psig range. This pressure will probably be needed to allow the oil to be dumped to a treater or tank and the water to be dumped to the water treating system. The higher the operating pressure, the smaller the compressor needed to compress the flash gas to sales. Compressor horsepower requirements are a function of the absolute discharge pressure divided by the absolute suction pressure.
Increasing the low-pressure separator pressure from 50 psig to 200 psig may decrease the compression horsepower required by 33%. However, it may also add backpressure to wells, restricting their flow, and allow more gas to be vented to atmosphere at the tank. Usually, an operating pressure of between 50 and 100 psig is optimum.
As stated before, the operating pressure of the highest-pressure separator will be no higher than the sales gas pressure. A possible exception to this could occur where the gas lift pressure is higher than the sales gas pressure. In choosing the operating pressures of the intermediate stages, it is useful to remember that the gas from these stages must be compressed.
Normally, this will be done in a multistage compressor. For practical reasons, the choice of separator operating pressures should match closely and be slightly greater than the compressor inter-stage pressures.
Fig. 3-10. Compressor stages and inlet points of separated gas from multistage separation.
The most efficient compressor sizing will be with a constant compressor ratio per stage. Therefore, an approximation of the intermediate separator operating pressures can be derived from
R = (Pd/Ps)1/n Eq. 3-2
where
R = Compression ratio per stage,
Pd = discharge pressure, psia,
Ps = suction pressure, psia,
n = number of stages.
In order to minimize inter-stage temperatures, the maximum ratio per stage will normally be in the range of 3.6 to 4.0. That means that most production facilities will have either two- or three-stage compressors. A two-stage compressor only allows for one possible intermediate separator operating pressure. A three-stage allows for either one operating at second- or third-stage suction pressure or two intermediate separators each operating at one of the two compressor intermediate suction pressures.( fig. 3-10).
Fundamentals of Oil and Gas Processing
Basics of Gas Field Processing
Basics of Corrosion in Oil and Gas Industry
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Basics of Gas Field Processing
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Re: Basics of Gas Field Processing Book "Full text"
Water Hydrocarbon Phase Behavior - Chapter 4
Chapter 4 70
Water Hydrocarbon Phase Behavior 70
4.1 Introduction 70
4.2 Measurements of Water Content of Gases 70
4.2.1 Bureau of Mines Dew-Point Tester (ASTM D 1142-63) 70
4.2.2 Electrolysis Method 72
4.2.3 Aluminum Oxide Humidity Sensor 74
4.2.4 Titration Method 74
4.2.5 Conductivity Cell 75
4.2.6 Dew-Point Tubes 75
4.2.7 Comparison of Methods 75
4.3 Water Content of Natural Gases 77
4.3.1 Water Content of Sweet Gases 77
4.3.2 Water Content of High CO2/H2S Gases 79
4.4 Prediction of Temperature drop due to pressure drop 81
4.5 Hydrates in Natural Gas Systems 82
4.5.1 Conditions which affect hydrate formation are: 85
4.5.2 Prediction of Sweet Natural Gas Hydrate Conditions 85
4.5.3 Hydrate Prediction Based on Composition for Sweet Gases 88
4.5.4 Hydrate Predictions for High CO2/H2S Content Gases 93
4.6 Hydrate Prevention 95
4.6.1 Adding Heat 95
4.6.2 Chemical Injection 96
4.6.3 Hydrate Inhibition with Methanol and Glycols 105
4.6.4 Low Dosage Hydrate Inhibitors (LDHIs) 110
-----------
Chapter 4
Water Hydrocarbon Phase Behavior
4.1 Introduction
As produced at the wellhead, natural gases are nearly always saturated with water. When water-saturated natural gas flows in a pipeline the following problems can occur:
1- Liquid water can collect in pipelines and so increase the pressure drop and/or cause slug flow.
2- Free water also can freeze into ice and/or form a solid hydrates and so reduce the gas flow or even plug the line completely.
3- Acid gases (H2S and CO2) dissolve in free water and can cause severe corrosion in internal surface of the pipeline.
Water removal and/or inhibition of hydrate formation is therefore a basic part of gas gathering. Design of oilfield gathering lines, dehydration, and hydrate inhibition facilities require two key phase-behavior predictions:
1- The water content of saturated natural gases
2- The hydrate formation temperate and pressures
4.2 Measurements of Water Content of Gases
Precautions required for gas sampling for dew point measurements are as follows:
1- The gas sample must be representative
2- The sampling line cannot contain free water
3- While flowing in the sample line, the gas must kept above its dew point temperature
4- If a glycol filter is installed, the sample must flow for five minutes (to saturate the filter with water) before the water content is measured.
The water content of natural gases can be measured by six different techniques: dew point, electrolysis, capacitance, conductivity, titration, and IR absorption. In addition dew-point tubes can provide approximate estimate. IR absorption is not used very often, and so the other methods are now summarized.
4.2.1 Bureau of Mines Dew-Point Tester (ASTM D 1142-63)
As shown in fig. 4-1, the Bureau of Mines dew-Point tester consists of a high-pressure, stainless steel or nickel-plated chamber. Gas entering through the inlet valve, A is directed by the deflector, B, onto the chilled, highly-polished stainless steel mirror, C, and then leaves the chamber through the outlet valve, D. The mirror, C, is cooled by cooling tube, F, which is attached to the chiller, G. Refrigerant enters the chiller at valve, H, and leaves at J Any dew on the mirror, C, is observed through the transparent Lucite window, E. The temperature at which dew is observed is read using the calibrated thermometer, K. Mirror M permits simultaneous viewing of the mirror C and reading of the thermometer, K.
Fig. 4-1. Bureau of Mines Dew-Point Tester
Dew-point measurement involves cooling the mirror, C, with a suitable refrigerant (C3 or Freon- 12 down to -20 0F, liquid CO2 down to -90 0F, dry ice/acetone down to –100 0F, or liquid N2 down to – 200 0F); flowing the sample gas over the polished mirror; and reading the temperature and pressure at which dew first appears and disappears on the mirror surface.
The following precautions will improve accuracy:
1- Use an illuminated magnifier and/or an LED temperature readout if the lighting makes it difficult to observe condensation.
2- Purge the tester to remove all air
3- Do not cool the mirror faster than 2 0F/min when within 5 0F dew point
4- While observing the mirror and thermometer, record the temperature at which dew first forms.
5- Let the mirror warm up and observe the temperature at which the dew disappears
6- Repeat steps 4 and 5 until the two temperatures agree within 2 0F
7- Take the average of the two temperatures as the dew point.
Liquid hydrocarbons, alcohols, or glycols also can condensate on the mirror before the water dew point is reached. The following characteristics distinguish water dew points. Refer to figure 4-2.
1- Water dew forms a distinct, opaque, grey circular spot in the center of the mirror (coldest spot). Water should not “wet” the mirror and should resist being blown off the mirror by increasing the gas flow. Ice crystals form an irregular white pattern against the previously –formed, grey water condensate. Barium sulfate and “water-cut” paste can confirm water dew point also.
2- In contrast, liquid-hydrocarbon condensates wet the mirror, expand in rainbow-like rings to cover all mirror, and can be “blown off” or “streak” the mirror by sudden increase in the sample gas flow rate.
3- Alcohol dew point appears as white spots with indistinct edges. Advanced alcohol spots are larger, increasingly white, and eventually form liquid drops that do not freeze.
4- Glycol dew points are darker, cover the entire mirror, and do not evaporate.
With the exception of the thermometer and pressure gauge, the Bureau of mines tester requires no calibration. The method is relatively inexpensive and easy to operate. However, this type of measurement can be time-consuming and cannot be recorded automatically. Accuracy can be very good but varies with the operator skill and dedication.
Fig. 4-2. Dew-point of water, hydrocarbons, and alcohols.
4.2.2 Electrolysis Method
The electrolysis method involves adsorbing and electrolyzing the water vapor in the sample gas. The heart of this instrument is an electrolytic cell consisting of the two 5-mil wires spirally wound throughout the inner wall of the insulating tube. A thin film of phosphorus pentoxide (P2O5) is applied between these two wires which are spaced 5 mil apart. As shown in figure 4-3, the sample gas first flows into the cell, then passes sensing windows covered with a semipermeable membrane, and finally exits. Water vapor, in direct proportional to the sample-gas concentration, is absorbed by the membrane, diffuses into the P2O5 film, and electrolyzed quantitatively. The resulting current is therefore, directly proportional to the water-vapor content of the sample gas.
The cell absorbs and electrolyzes moisture at fractional parts-per-million (ppm) or other units of measure. How: One hundred percent of the sample moisture is absorbed by a phosphorus pentoxide (P2O5) film that covers two spirally-wound electrodes embedded in a hollow glass tube. When the sample gas enters the cell at a known flow rate, the film absorbs all the moisture molecules present. By applying an electrical potential (voltage) to the electrodes, each absorbed water molecule is electrolyzed, generating a finite current. This current is precise and proportional to the amount of absorbed water. It is, therefore, an exact, direct measurement of the water vapor present in the sample gas.
Only the ammeter and flowmeter require calibration, and this makes the electrolytic method one of the most accurate and fundamental available. Ammeter calibration requires “calibration gas” (Natural gas with a known water content). Moisture calibrators provide a continuous supply of calibration gas by saturating the gas with water at 32 0F. Water content from 9.5 to 170 lb/MMscf, can be obtained by varying the calibration gas pressure.
Note the following sources of errors:
1- Contamination of the cell or coating of the P2O5 strip by oil, condensate, glycol, compressor oil, etc. (anything that changes the adsorption of the water vapor). Such a contaminated cell will exhibit a low reading (0.25 lb H2O/MMscf) that will not change when the sample gas flow-rate is varied.
2- Washout of the cell by excess water, alcohol, oil, methane, amine, etc., produces an essentially zero meter reading.
3- A dead short produces an off-scale meter reading.
Models without a semipermeable membrane to protect the P2O5 membrane are far more susceptible to contamination and washout. Two additional warning are worthwhile:
1- The electrolytic cell does not operate well below 32 0F, and should be temperature controlled if necessary.
2- Phosphoric acid can cause severe harm to skin and eyes. Extreme caution should be exercised when the electrolytic cell is cleaned and recoated.
Fig. 4-3. Electrolysis Method.
4.2.3 Aluminum Oxide Humidity Sensor
The moisture sensor consists of a thin, porous layer or film of aluminum oxide (AL2O3) sandwiched between two electrodes. Refere to figure 4-4.
The sandwich sensor is essentially a capacitor, with the AL2O3 the dielectric. When an AC voltage is applied, the resulting impedance varies with the amount of water absorbed in the aluminum oxide film. In turn, the quantity of adsorbed water depends on the partial pressure of the water vapor in the gas sample flowing around the sensor. A suitable electronic circuit converts the measured impedance to the desired units of water vapor content. The sensor or probe is built so that the water vapor equilibrates rapidly.
This capacitance method is used to measure water dew point ranging from -150 to 70 0F with a response time of less than 5 seconds for a 63% step change in moisture content. As with the electrolysis method, contamination by pipe-scale, carbon, salt, and conductive liquids (glycol, methanol) can impede measurement. The sensor is not harmed by liquid slugs of condensate, methanol, glycol, and water. Proper cleaning restores the sensor.
4.2.4 Titration Method
The water content is determined by titration using a water specific reagent (usually Karl Fischer reagent). The sample gas enters the reaction cell, bubbles through a small known quantity of liquid reagent (0.5 mL), and exits via a reagent trap, a pressure reducing regulator, and finally a flow meter. A pair of platinum electrodes, sense the end point of the titration (when the entering water vapor has exhausted the batch of liquid reagent). Then a fresh batch of reagent is injected into the reaction cell by pump.
Electronic circuitry measures the time between end points and the sample gas flow-rate and pressure then it computes and then displays the water content. The entire cycle takes about two seconds.
Karl Fischer reagent is inert to hydrocarbons, carbon oxide, glycol, amines, and most sulfur compounds, e.g., odorizing mercaptans. Hydrogen sulfide will cause the moisture titrator to read high by 0.7 lb H2O/MMscf per grain H2S/100 scf.
Fig. 4-4. Aluminum oxide humidity sensor.
Fig. 4-5. Karl Fischer titrator for moisture content in gas.
4.2.5 Conductivity Cell
The Hygromat measuring cell consists of two stainless steel plates separated and electrically insulated from each other by a ceramic layer. The ceramic layer has eight holes which are partially filled with a hygroscopic salt-glycerol solution. Water is absorbed reversibly by the hygroscopic solution until equilibrium is reached with the surrounding natural gas. In turn the conductivity of the salt-glycerol solution increases as water is absorbed and decreases when water is desorbed.
4.2.6 Dew-Point Tubes
The dew-point tube uses a sampling pump as in figure 4-6. A water detector tube is placed in the pump and 100 mL of pipeline gas is pulled through the tube. The detector tube figure 4-7, is filled with magnesium perchlorate contained in a fine silica gel. Water vapor is absorbed by magnesium perchlorate to produce an alkaline reaction that changes the color of the Hammet’s indicator (crystal violet). The water content in lb/MMscf is read directly from the length of the stain in the detector tube. Overall accuracy is +/- 25%. Alcohols, glycols, and amines cause high readings. The tube range is usually 6-80 lb H2O/MMscf.
4.2.7 Comparison of Methods
- The Bureau of mines dew point method is respected as the one defined by an ASTM standard. Equipment cost is low, but the method is labor intensive in that a single measurements requires approximately 15 minutes. Accuracy varies with operator skill. Uncertainties of +/- 1 0F are attainable for dew points above 32 0F but increase to +/- 4 0F for dew points between -80 to -100 0F.
- The popularity of the electrolytic moisture analyzers is increasing significantly. Their accuracy compares favorably with the dew-point method. Advantages include light weight, probability, continuous readings, fast response times, and ready interfacing with alarms and other process monitors. Improved methods of cleaning and recoating the P2O5 film reduce the most frequent disadvantages of contamination and washout.
- The aluminum oxide sensor is relatively recent and exhibits accuracies and advantages similar to the electrolytic analyzer. It is especially suitable for every dry gases. However, response time is slow and removal of contaminants is more difficult.
- The titrator equipment is relatively expensive but not readily portable. The main advantages are accuracy (3% of reading) and immunity to contaminants, such as glycols and alcohols. The chief disadvantage is the hazardous nature of the Karl Fischer reagents, which creates a disposal problem.
- The advantages of the conductivity method include long-term stability. The 63% response time to sudden changes in gas humidity varies from 5 to 30 minutes depending on gas flow rate and pressure. One disadvantage is that conductivity varies with temperature and so the meter must be kept at constant temperature.
- Detector tubes provide inexpensive approximate estimate. They can be used by nontechnical personnel with minimum training.
Fig. 4-6. Sampling equipment and methods for dew-point tubes.
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Fig. 4-7. Dew-point tube readings.
4.3 Water Content of Natural Gases
The saturated water content of a gas depends on pressure, temperature, and composition. The effect of composition increases with pressure and is particularly important if the gas contains CO2 and/or H2S. For lean, sweet natural gases containing over 70% methane and small amounts of heavy hydrocarbons, generalized pressure-temperature correlations are suitable for many applications.
4.3.1 Water Content of Sweet Gases
Fig. 4-8 is an example of one such correlation which has been widely used for many years in the design of “sweet” natural gas dehydrators. The gas gravity correlation should never be used to account for the presence of H2S and CO2 and may not always be adequate for certain hydrocarbon effects, especially for the prediction of water content at pressures above 1500 psia. The hydrate formation line is approximate and should not be used to predict hydrate formation conditions.
Example 4-1 — Determine the saturated water content for a sweet lean hydrocarbon gas at 150°F and 1,000 psia.
From Fig. 4-8, W = 220 lb/MMscf.
For a 26 molecular weight gas, Cg = 0.98 (Correction for gas gravity- fig. 4-8)
W = (0.98)(220) = 216 lb/MMscf
For a gas in equilibrium with a 3% brine, Cs = 0.93 (Correction for salinity fig. 4-8)
W = (0.93)(220) = 205 lb/MMscf
Fig. 4-8. Water Content of sweet Hydrocarbon Gas. McKetta and Wehe, 1958; GPSA, 1987.
4.3.2 Water Content of High CO2/H2S Gases
Saturated water content of pure CO2 and H2S can be significantly higher than that of sweet natural gas, particularly at pressures above about 700 psia at ambient temperatures.
Corrections for H2S and CO2 should be applied when the gas mixture contains more than 5% H2S and/or CO2 at pressures above 700 psia. These corrections become increasingly significant at higher concentrations and higher pressures.
Below 40% acid gas components, one method of estimating the water content uses Eq 4-1 and Fig. 4-8, 4-9, and 4-10.
W = yHC WHC + yCO2 WCO2 + yH2SWH2S Eq 4-1
Where
W = Water content of gas, lb/MMscf
yHC = Mole fraction of hydrocarbon in the gas phase
WHC = Water content in hydrocarbon gas, lb/MMscf
YCO2 = Mole fraction of CO2 in the gas phase
WCO2 = Effective water content in CO2 gas, lb/MMscf
yH2S = Mole fraction of H2S in the gas phase
WH2S= Effective water content in H2S gas, lb/MMscf
Fig. 4-9. Effective Water Content of H2S in Natural Gas Mixtures vs. Temperature at Various Pressures
Fig. 4-10. Effective Water Content of CO2 in Natural Gas Mixtures vs. Temperature at Various Pressures
Note that Fig. 4-9 and 4-10 provide values for what is termed the “effective” water content of CO2 and H2S in natural gas mixtures for use only in Eq 4-1. These are not pure CO2 and H2S water contents.
A second method is Fig. 4-11. The CO2 is converted to equivalent H2S, using the factor 70%.
Example 4-2 — Determine the saturated water content of an 80% C1, 20% CO2 mixture at 160 °F and 2000 psia. The experimentally determined water content was 172 lb/MMscf.
Method One
WHC = 167 lb/MMscf (Fig. 4-8)
WCO2 = 240 lb/MMscf (Fig. 4-10)
W = (0.80 x 167) + (0.20 x 240)
=182 lb/MMscf
Method Two
First the composition must be converted for use with Fig. 4-11.
yH2S (pseudo) = 0.70 (yCO2 ) = 0.70 (0.20) = 0.14
Enter the left side of Fig. 4-11 at 160°F and move to the % H2S Equivalent line (14%). Proceed vertically upward to the Pressure, psia line (2000 psia), and move horizontally to the left to Water Content Ratio scale (ratio of 1.16).
W = (1.16)(167) = 194 lb/MMscf
Fig.4-11 . Calculated Water Content of Acid Gas Mixtures.
4.4 Prediction of Temperature drop due to pressure drop
Figure 4-12 can be used to get a quick approximate solution for the temperature drop of a natural gas stream (accuracy is +/- 5%.). For example, if the initial pressure is 4,000 psi and the final pressure is 1,000 psi, P is 3,000 psi, the change in temperature is 80°F. This curve is based on a liquid concentration of 20 bbl/MMscf. The greater the amount of liquid in the gas the lower the temperature drop, that is, the higher the calculated final temperature. For each increment of 10 bbl/MMscf there is a correction of 5°F. For example, if there is no liquid, the final temperature is 10°F cooler (the temperature drop is 10°F more) than indicated by Figure 4-12.
Example 4-3: Determine the temperature drop across a choke
Given: A well with a flowing tubing pressure of 4000 psi and 20 bbl of hydrocarbon condensate and a downstream back pressure of 1000 psi.
Solution: Initial pressure = 4000 psi
Final pressure = 1000 psi , P = 3000 psi
From Figure 4-12 correlation; intersect initial pressure = 4000 and ΔP-3000 read ΔT = 800F.
4.5 Hydrates in Natural Gas Systems
A hydrate is a physical combination of water and other small molecules to produce a solid which has an “ice-like” appearance but possesses a different structure than ice. Their formation in gas and/or NGL systems can plug pipelines, equipment, and instruments, restricting or interrupting flow.
There are three recognized crystalline structures for such hydrates. Where, water molecules build the lattice and hydrocarbons, nitrogen, CO2 and H2S occupy the cavities. Smaller molecules (CH4, C2H6, CO2, H2S) stabilize a body-centered cubic called Structure I. Larger molecules (C3H8, i-C4H10, n-C4H10) form a diamond-lattice called Structure II.
Normal paraffin molecules larger than n-C4H10 do not form Structure I and II hydrates as they are too large to stabilize the lattice. However, some isoparaffins and cycloalkanes larger than pentane are known to form Structure H hydrates.
Gas composition determines structure type. Mixed gases will typically form Structure II. From a practical viewpoint, the structure type does not affect the appearance, properties, or problems caused by the hydrate. It does however, have a significant effect on the pressure and temperature at which hydrates form. Structure II hydrates are more stable than Structure I. This is why gases containing C3H8 and i-C4H10 will form hydrates at higher temperatures than similar gas mixtures which do not contain these components. The effect of C3H8 and i-C4H10 on hydrate formation conditions can be seen in Fig. 4-15. At 1000 psia, a 0.6 sp. gr. gas (composition is shown in Fig. 4-15) has a hydrate formation temperature which is 12-13°F higher than pure methane.
Fig.4-13. Hydrate structures, typical hydrate plugging, and illustration of hydrate formation.
Fig. 4-14. Conditions for Hydrate Formation for Light Gases
Fig. 4-15. Pressure-Temperature Curves for Predicting Hydrate Formation. Katz.1945.GPSA 1987.
The presence of H2S in natural gas mixtures results in a substantially warmer hydrate formation temperature at a given pressure. CO2, in general, has a much smaller impact and often reduces the hydrate formation temperature at fixed pressure for a hydrocarbon gas mixture.
4.5.1 Conditions which affect hydrate formation are:
Primary Considerations
• Gas or liquid must be at or below its water dew point or saturation condition (Note: liquid water does not have to be present for hydrates to form)
• Temperature
• Pressure
• Composition
Secondary Considerations
• Mixing
• Kinetics
• Physical site for crystal formation and agglomeration such as a pipe elbow, orifice, thermowell, or line scale
• Salinity
In general, hydrate formation will occur as pressure increases and/or temperature decreases to the hydrate formation condition.
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4.5.2 Prediction of Sweet Natural Gas Hydrate Conditions
Fig. 4-14, based on experimental data, presents the hydrate pressure-temperature equilibrium curves for pure methane, ethane, propane, and for a nominal 70% ethane 30% propane mix.
Fig. 4-15 through 4-19, based on gas gravity, may be used for first approximations of hydrate formation conditions and for estimating permissible expansion of sweet natural gases without the formation of hydrates.
The conditions at which hydrates can form are strongly affected by gas composition.
Example 4-4- Find the pressure at which hydrate forms for a gas with the following composition. T = 50°F.
Table 4-1. Gas composition and molecular weight calculation for Example 4-4.
Gas specific gravity = MWgas/MWair = 20.08/28.964 = 0.693
From Fig. 4-15 at 50°F.
P = 320 psia for 0.7 gravity gas
Example 4-5— The gas in Example 4-4 is to be expanded from 1,500 psia to 500 psia. What is the minimum initial temperature that will permit the expansion without hydrate formation?
The 1,500 psia initial pressure line and the 500 psia final pressure line intersect just above the 110°F curve on Fig. 4-17. Approximately 112°F is the minimum initial temperature.
Example 4-6 — How far may a 0.6 gravity gas at 2,000 psia and 100°F be expanded without hydrate formation?
On Fig. 4-16 find the intersection of 2,000 initial pressure line with the 100°F initial temperature curve. Read on the x-axis the permissible final pressure of 1100 psia.
Fig. 4- 16. Permissible Expansion of a 0.6-Gravity Natural Gas Without Hydrate Formation
Example 4-7 — How far may a 0.6 gravity gas at 2,000 psia and 140°F be expanded without hydrate formation?
On Fig. 4-16, the 140°F initial temperature curve does not intersect the 2,000 psia initial pressure line. Therefore, the gas may be expanded to atmospheric pressure without hydrate formation.
Conditions predicted by Fig. 4-15 through 4-19 may be significantly in error for compositions other than those used to derive the charts. For more accurate determination of hydrate formation conditions, the following procedures should be followed. In addition, fig. 4-15 through 4-19, do not account for liquid water and liquid hydrocarbons present or formed during the expansion. These can have a significant effect on the outlet temperature from the pressure reduction device.
Fig. 4- 19. Permissible Expansion of a 0.9-Gravity Natural Gas Without Hydrate Formation
4.5.3 Hydrate Prediction Based on Composition for Sweet Gases
1- Katz’s Graph
The first method is for approximate and fast prediction, using of Katz graph represented in figure 4-15. On the other hand, for below 1,000 psi (70 bar), the figure can be approximated by:
t(°F) = −16.5 – [6.83/(SpGr)2] + 13.8 ln[P(psia)] Eq 4-9.
2- Katz’s Vapor solid equilibrium method
Several correlations have proven useful for predicting hydrate formation of sweet gases and gases containing minimal amounts of CO2 and/or H2S. The most reliable ones require a gas analysis. The Katz method utilizes vapor solid equilibrium constants defined by the Eq 4-10.
Kvs = y / xs Eq 4-10
Where
Kvs = vapor/solid equilibrium K-value
y = mole fraction in the gas phase
xs = mole fraction in the solid phase
WARNING: Not good for pure components – only mixtures.
The Katz’s correlation is not recommended above 1000-1500 psia, depending on composition.
The applicable K-value correlations for the hydrate forming molecules (methane, ethane, propane, isobutane, normal butane, carbon dioxide, and hydrogen sulfide) are shown in Fig. 4-20 to 4-26. Normal butane cannot form a hydrate by itself but can contribute to hydrate formation in a mixture.
For calculation purposes, all molecules too large to form hydrates have a K-value of infinity. These include all normal paraffin hydrocarbon molecules larger than normal butane.
Nitrogen is assumed to be a non-hydrate former and is also assigned a K-value of infinity.
The Kvs values are used in a “dewpoint” equation to determine the hydrate temperature or pressure. The calculation is iterative and convergence is achieved when the following objective function (Eq 4-3) is satisfied.
∑y/Kvs = 1.0 Eq. 4-11
Fig. 4- 20. Kvs. Vapor solid equilibrium constants for Methane.
Fig. 4- 21. Kvs. Vapor solid equilibrium constants for Ethane.
Fig. 4- 22. Kvs. Vapor solid equilibrium constants for Propane.
Fig. 4- 23. Kvs. Vapor solid equilibrium constants for Iso-Butane.
Fig. 4- 24. Kvs. Vapor solid equilibrium constants for N-Butane.
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Fig. 4- 25. Kvs. Vapor solid equilibrium constants for carbon dioxide.
Fig. 4- 26. Kvs. Vapor solid equilibrium constants for hydrogen sulfide.
Example 4-8 — Calculate the pressure for hydrate formation at 50°F for a gas with the following composition.
Table. 4-2. Example. 4-8.
The ∑y/Kvs value is slightly over than 300 psia, by iterpolating linearly, = 1.0 @ 305 psia.
Hydrate pressure at temperature 50 0F = 305 psia
Third and fifth columns are values obtained from chart for each component at temperature and pressure values.
Fourth column contains the results of dividing Mole fraction of gas by third column.
Sixth column contains the results of dividing Mole fraction of gas by fifth column.
3- Motiee (1991) suggested the following equation for hydrate temperature prediction:
T(0F) = -238.24469 + 78.99667 log P(psi) – 5.352544 (log P(psi) )2 +
349.473877 y – 150.854675 y2 – 27.604065 y log P(psi) Eq. 4-12
Where ᵞ is gas specific gravity.
This equation is well known and widely used in the oil and gas industry because of its accuracy for natural gas mixtures.
4- Recently (2009), a new correlation developed by Bahadori and Vuthaluru, with specific gravities from 0.55 to 1, shows the best efficiency. The equation is suitable for estimating the HFT, especially for natural gas mixtures:
T(k) = AyB (ln P(KPa))C Eq. 4-13
where:
A=194.681789
B=0.044232
C=0.189829
0F = (0K - 273.15) x 9/5 + 32
0K = (0F - 32) x 5/9 + 273.15
(0C=0K-273.15)
psi = 6.895 KPa,
4.5.4 Hydrate Predictions for High CO2/H2S Content Gases
The Katz method of predicting hydrate formation temperature gives reasonable results for sweet paraffin hydrocarbon gases. The Katz method should not be used for gases containing significant quantities of CO2 and/or H2S despite the fact that Kvs values are available for these components. Hydrate formation conditions for high CO2/H2S gases can vary significantly from those composed only of hydrocarbons. The addition of H2S to a sweet natural gas mixture will generally increase the hydrate formation temperature at a fixed pressure. A method by Baille & Wichert for predicting the temperature of high H2S content gases is shown in Fig. 4-27.
Fig.4-27. Hydrate Chart for Gases Containing H2S
4.6 Hydrate Prevention
Hydrate prevention is accomplished by keeping the:
1- Operating conditions must remain out of the hydrate-formation zone by heating or temperature control.
2- Hydrate point must be maintained below the operating conditions of the system by chemical treatment.
Two common methods of hydrate-formation prevention are:
1- Temperature control
2- Chemical injection
Example 4-9 — Estimate the hydrate formation temperature at 610 psia of a gas with the following analysis using Fig. 4-27.
Table. 4-3. Solution of example 4-9.
MW = 19.75
Sp.Gr (ᵧ) = 0.682
Solution Steps:
1. Enter left side of Fig. 4-27 at 600 psia and proceed to the H2S concentration line (4.18 mol%)
2. Proceed down vertically to the specific gravity of the gas (ᵞ = 0.682)
3. Follow the diagonal guide line to the temperature at the bottom of the graph (T = 63.5°F).
4. Apply the C3 correction using the insert at the upper left.
Enter the left hand side at the H2S concentration and proceed to the C3 concentration line (0.67%). Proceed down vertically to the system pressure and read the correction on the left hand scale (–2.7°F)
Note: The C3 temperature correction is negative when on the left hand side of the graph and positive on the right hand side.
TH = 63.5 −2.7 = 60.8°F
Fig. 4-27 was developed based on calculated hydrate conditions using the Peng-Robinson EOS. It has proven quite accurate when compared to the limited amount of experimental Mole Fraction data available. It should only be extrapolated beyond the experimental data base with caution.
4.6.1 Adding Heat
Adding heat is effective because hydrates normally do not occur above 70 0F.
It offers a simple and economical solution for land and offshore facilities (if waste heat is available).
Flow stream is preheated, either through an indirect line heater or heat exchanger, before passing through a choke. Flow stream is then reheated to maintain the temperature above the hydrate formation temperature.
A major drawback in offshore installations is that it is almost impossible to maintain flowline temperatures significantly above the water temperature if the flowlines extend more than a few hundred feet under water. Thus, either the “free water” must be separated while still at temperature or an alternate method
selected.
4.6.1.1 Temperature Control
Wellhead Indirect Heaters
An indirect heater is used to heat gas to maintain temperatures above that of the hydrate formation.
It consists of an atmospheric vessel containing a fire tube (usually fired by gas, steam, or heating oil) and a coil (designed to withstand shut in tubing pressure “SITP”) that is heated by the intermediate fluid (usually water) and the fluid is heated. The fire tube and coil are immersed in a heat transfer fluid (normally water), and heat is transferred to the fluid in the coil.
Figure 4-28 shows a typical heater installation at the wellhead.
Long-Nose Heater Choke (Figure 4-29).
A long body choke installed in the indirect heater to position the choke orifice within the indirect heater bath. Since the walls of the choke orifice are heated by the water bath, hydrates will not form in the orifice and cause plugging.
Flowline indirect Heaters
Flowline heaters differ from wellhead heaters in purpose only.
The purpose of a wellhead heater is to heat the flow stream at or near the wellhead where choking or pressure reduction occurs.
The purpose of a flowline heater is to provide additional heat if required.
The design is the same as an indirect heater except that the choke, shut-in, and relief equipment are seldom used.
System Optimization
System operation has to be optimized before heaters can be effectively designed and located.
Heat requirements that appear to be large can often be reduced to minimal values or even eliminated by revising the mode of operation. For example: Fields having multiple producing wells can be combined to use higher flowing temperatures thus minimizing heater requirements.
If reducing the gas stream pressure is necessary, it is generally more efficient to do so at a central point where the necessary heater fuel gas can be obtained from separators or scrubbers.
Requires flowline wall thickness to be increased so as to withstand wellhead SITP.
An alternative is to install wellhead shut-down valves and flowline high pressure switches.
Downhole Regulators
Downhole regulators are feasible for high capacity gas wells at locations where certain risks to other downhole equipment are acceptable. The theory behind the use of a downhole regulator is that the pressure drop from flowing pressure to near-sales line pressure is taken downhole where the formation temperature is sufficient to prevent hydrate formation. The tubing string above the regulator then acts as a subsurface heater. Calculations involved in downhole regulator design are rather involved. They depend on characteristics such as:
Wellbore configuration, flowing downhole pressures and temperature, and well depth.
Although shortcut procedures are available to estimate the feasibility of downhole regulators, tool company representatives can provide detailed design information.
4.6.2 Chemical Injection
The formation of hydrates can also be prevented by dehydrating the gas or liquid to eliminate the formation of a condensed water (liquid or solid) phase. In some cases, however, dehydration may not be practical or economically feasible. In these cases, chemical inhibition can be an effective method of preventing hydrate formation. Chemical inhibition utilizes injection of thermodynamic inhibitors or low dosage hydrate inhibitors (LDHIs). Thermodynamic inhibitors are the traditional inhibitors (i.e., one of the glycols or methanol), which lower the temperature of hydrate formation. LDHIs are either kinetic hydrate inhibitors (KHIs) or antiagglomerants (AAs).
They do not lower the temperature of hydrate formation, but do diminish its effect. KHIs lower the rate of hydrate formation, which inhibits its development for a defined duration.
AAs allow the formation of hydrate crystals but restrict them to sub-millimeter size.
Fig. 4-28. Wellhead indirect heater schematic.
4.6.2.1 Thermodynamic Inhibitors
Inhibition utilizes injection of one of the glycols or methanol into a process stream where it can combine with the condensed aqueous phase to lower the hydrate formation temperature at a given pressure.
Both glycol and methanol can be recovered with the aqueous phase, regenerated and re-injected. For continuous injection in services down to –40°F, one of the glycols usually offers an economic advantage versus methanol recovered by distillation.
At cryogenic conditions (below –40°F) methanol usually is preferred because glycol’s viscosity makes effective separation difficult.
Ethylene glycol (EG), diethylene glycol (DEG), and triethylene glycol (TEG), glycols have been used for hydrate inhibition. The most popular has been ethylene glycol because of its lower cost, lower viscosity, and lower solubility in liquid hydrocarbons.
Freezing point of aqueous methanol solution is given in fig. 4-30.
Hydrate inhibitors are used to lower the hydrate formation temperature of the gas.
Recovery and regeneration steps are used in all continuous glycol injection projects and in several large-capacity methanol injection units.
Injection of hydrate inhibitors should be considered for the following applications:
• Pipeline systems in which hydrate trouble is of short duration
• Gas pipelines that operate at a few degrees below the hydrate formation temperature
• Gas-gathering systems in pressure-declining fields
• Gas lines in which hydrates form as localized points
Methanol and the lower molecular weight glycols have the most desirable characteristics for use as hydrate inhibitors.
When hydrate inhibitors are injected in gas flowlines or gathering systems, installation of a free-water knockout (FWKO) at the wellhead proves to be economical in nearly every case.
Removing the free water from the gas steam reduces the amount of inhibitor required.
To be effective, the inhibitor must be present at the very point where the wet gas is cooled to its hydrate temperature. For example, in refrigeration plants glycol inhibitors are typically sprayed on the tube-sheet faces of the gas exchangers so that it can flow with the gas through the tubes. As water condenses, the inhibitor is present to mix with the water and prevent hydrates. Injection must be in a manner to allow good distribution to every tube or plate pass in chillers and heat exchangers operating below the gas hydrate temperature.
Table 4-4 lists some physical properties of methanol and the lower molecular weight glycols.
The inhibitor and condensed water mixture is separated from the gas stream along with a separate liquid hydrocarbon stream. At this point, the water dewpoint of the gas stream is essentially equal to the separation temperature. Glycol-water solutions and liquid hydrocarbons can emulsify when agitated or when expanded from a high pressure to a lower pressure, e.g., JT expansion valve. Careful separator design will allow nearly complete recovery of the diluted glycol for regeneration and reinjection.
Fig. 4-29. Indirect heater and long-Nose choke.
Fig. 4-30. Freezing Points of Aqueous Methanol Solutions
Methanol Injection Considerations
Methanol is well-suited for use as a hydrate inhibitor because it is:
Noncorrosive
Nonreactive chemically with any constituent of the gas
Soluble in all proportions in water
Volatile under pipeline conditions
Reasonable in cost
Of a vapor pressure greater than that of water
Methanol is injected by means of an injection pump (3 in Figure 4-31) into the flowline upstream of the choke or pressure control valve (2). A temperature controller may be installed to measure the temperature of the gas in the low-pressure flowlineand adjusts the methanol rate accordingly.
Fig. 4-31. Methanol injection system at wellhead.
------------Fundamentals of Oil and Gas Processing Book
Basics of Gas Field Processing Book
Prediction and Inhibition of Gas Hydrates Book
Basics of Corrosion in Oil and Gas Industry Book
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Chapter 4 70
Water Hydrocarbon Phase Behavior 70
4.1 Introduction 70
4.2 Measurements of Water Content of Gases 70
4.2.1 Bureau of Mines Dew-Point Tester (ASTM D 1142-63) 70
4.2.2 Electrolysis Method 72
4.2.3 Aluminum Oxide Humidity Sensor 74
4.2.4 Titration Method 74
4.2.5 Conductivity Cell 75
4.2.6 Dew-Point Tubes 75
4.2.7 Comparison of Methods 75
4.3 Water Content of Natural Gases 77
4.3.1 Water Content of Sweet Gases 77
4.3.2 Water Content of High CO2/H2S Gases 79
4.4 Prediction of Temperature drop due to pressure drop 81
4.5 Hydrates in Natural Gas Systems 82
4.5.1 Conditions which affect hydrate formation are: 85
4.5.2 Prediction of Sweet Natural Gas Hydrate Conditions 85
4.5.3 Hydrate Prediction Based on Composition for Sweet Gases 88
4.5.4 Hydrate Predictions for High CO2/H2S Content Gases 93
4.6 Hydrate Prevention 95
4.6.1 Adding Heat 95
4.6.2 Chemical Injection 96
4.6.3 Hydrate Inhibition with Methanol and Glycols 105
4.6.4 Low Dosage Hydrate Inhibitors (LDHIs) 110
-----------
Chapter 4
Water Hydrocarbon Phase Behavior
4.1 Introduction
As produced at the wellhead, natural gases are nearly always saturated with water. When water-saturated natural gas flows in a pipeline the following problems can occur:
1- Liquid water can collect in pipelines and so increase the pressure drop and/or cause slug flow.
2- Free water also can freeze into ice and/or form a solid hydrates and so reduce the gas flow or even plug the line completely.
3- Acid gases (H2S and CO2) dissolve in free water and can cause severe corrosion in internal surface of the pipeline.
Water removal and/or inhibition of hydrate formation is therefore a basic part of gas gathering. Design of oilfield gathering lines, dehydration, and hydrate inhibition facilities require two key phase-behavior predictions:
1- The water content of saturated natural gases
2- The hydrate formation temperate and pressures
4.2 Measurements of Water Content of Gases
Precautions required for gas sampling for dew point measurements are as follows:
1- The gas sample must be representative
2- The sampling line cannot contain free water
3- While flowing in the sample line, the gas must kept above its dew point temperature
4- If a glycol filter is installed, the sample must flow for five minutes (to saturate the filter with water) before the water content is measured.
The water content of natural gases can be measured by six different techniques: dew point, electrolysis, capacitance, conductivity, titration, and IR absorption. In addition dew-point tubes can provide approximate estimate. IR absorption is not used very often, and so the other methods are now summarized.
4.2.1 Bureau of Mines Dew-Point Tester (ASTM D 1142-63)
As shown in fig. 4-1, the Bureau of Mines dew-Point tester consists of a high-pressure, stainless steel or nickel-plated chamber. Gas entering through the inlet valve, A is directed by the deflector, B, onto the chilled, highly-polished stainless steel mirror, C, and then leaves the chamber through the outlet valve, D. The mirror, C, is cooled by cooling tube, F, which is attached to the chiller, G. Refrigerant enters the chiller at valve, H, and leaves at J Any dew on the mirror, C, is observed through the transparent Lucite window, E. The temperature at which dew is observed is read using the calibrated thermometer, K. Mirror M permits simultaneous viewing of the mirror C and reading of the thermometer, K.
Fig. 4-1. Bureau of Mines Dew-Point Tester
Dew-point measurement involves cooling the mirror, C, with a suitable refrigerant (C3 or Freon- 12 down to -20 0F, liquid CO2 down to -90 0F, dry ice/acetone down to –100 0F, or liquid N2 down to – 200 0F); flowing the sample gas over the polished mirror; and reading the temperature and pressure at which dew first appears and disappears on the mirror surface.
The following precautions will improve accuracy:
1- Use an illuminated magnifier and/or an LED temperature readout if the lighting makes it difficult to observe condensation.
2- Purge the tester to remove all air
3- Do not cool the mirror faster than 2 0F/min when within 5 0F dew point
4- While observing the mirror and thermometer, record the temperature at which dew first forms.
5- Let the mirror warm up and observe the temperature at which the dew disappears
6- Repeat steps 4 and 5 until the two temperatures agree within 2 0F
7- Take the average of the two temperatures as the dew point.
Liquid hydrocarbons, alcohols, or glycols also can condensate on the mirror before the water dew point is reached. The following characteristics distinguish water dew points. Refer to figure 4-2.
1- Water dew forms a distinct, opaque, grey circular spot in the center of the mirror (coldest spot). Water should not “wet” the mirror and should resist being blown off the mirror by increasing the gas flow. Ice crystals form an irregular white pattern against the previously –formed, grey water condensate. Barium sulfate and “water-cut” paste can confirm water dew point also.
2- In contrast, liquid-hydrocarbon condensates wet the mirror, expand in rainbow-like rings to cover all mirror, and can be “blown off” or “streak” the mirror by sudden increase in the sample gas flow rate.
3- Alcohol dew point appears as white spots with indistinct edges. Advanced alcohol spots are larger, increasingly white, and eventually form liquid drops that do not freeze.
4- Glycol dew points are darker, cover the entire mirror, and do not evaporate.
With the exception of the thermometer and pressure gauge, the Bureau of mines tester requires no calibration. The method is relatively inexpensive and easy to operate. However, this type of measurement can be time-consuming and cannot be recorded automatically. Accuracy can be very good but varies with the operator skill and dedication.
Fig. 4-2. Dew-point of water, hydrocarbons, and alcohols.
4.2.2 Electrolysis Method
The electrolysis method involves adsorbing and electrolyzing the water vapor in the sample gas. The heart of this instrument is an electrolytic cell consisting of the two 5-mil wires spirally wound throughout the inner wall of the insulating tube. A thin film of phosphorus pentoxide (P2O5) is applied between these two wires which are spaced 5 mil apart. As shown in figure 4-3, the sample gas first flows into the cell, then passes sensing windows covered with a semipermeable membrane, and finally exits. Water vapor, in direct proportional to the sample-gas concentration, is absorbed by the membrane, diffuses into the P2O5 film, and electrolyzed quantitatively. The resulting current is therefore, directly proportional to the water-vapor content of the sample gas.
The cell absorbs and electrolyzes moisture at fractional parts-per-million (ppm) or other units of measure. How: One hundred percent of the sample moisture is absorbed by a phosphorus pentoxide (P2O5) film that covers two spirally-wound electrodes embedded in a hollow glass tube. When the sample gas enters the cell at a known flow rate, the film absorbs all the moisture molecules present. By applying an electrical potential (voltage) to the electrodes, each absorbed water molecule is electrolyzed, generating a finite current. This current is precise and proportional to the amount of absorbed water. It is, therefore, an exact, direct measurement of the water vapor present in the sample gas.
Only the ammeter and flowmeter require calibration, and this makes the electrolytic method one of the most accurate and fundamental available. Ammeter calibration requires “calibration gas” (Natural gas with a known water content). Moisture calibrators provide a continuous supply of calibration gas by saturating the gas with water at 32 0F. Water content from 9.5 to 170 lb/MMscf, can be obtained by varying the calibration gas pressure.
Note the following sources of errors:
1- Contamination of the cell or coating of the P2O5 strip by oil, condensate, glycol, compressor oil, etc. (anything that changes the adsorption of the water vapor). Such a contaminated cell will exhibit a low reading (0.25 lb H2O/MMscf) that will not change when the sample gas flow-rate is varied.
2- Washout of the cell by excess water, alcohol, oil, methane, amine, etc., produces an essentially zero meter reading.
3- A dead short produces an off-scale meter reading.
Models without a semipermeable membrane to protect the P2O5 membrane are far more susceptible to contamination and washout. Two additional warning are worthwhile:
1- The electrolytic cell does not operate well below 32 0F, and should be temperature controlled if necessary.
2- Phosphoric acid can cause severe harm to skin and eyes. Extreme caution should be exercised when the electrolytic cell is cleaned and recoated.
Fig. 4-3. Electrolysis Method.
4.2.3 Aluminum Oxide Humidity Sensor
The moisture sensor consists of a thin, porous layer or film of aluminum oxide (AL2O3) sandwiched between two electrodes. Refere to figure 4-4.
The sandwich sensor is essentially a capacitor, with the AL2O3 the dielectric. When an AC voltage is applied, the resulting impedance varies with the amount of water absorbed in the aluminum oxide film. In turn, the quantity of adsorbed water depends on the partial pressure of the water vapor in the gas sample flowing around the sensor. A suitable electronic circuit converts the measured impedance to the desired units of water vapor content. The sensor or probe is built so that the water vapor equilibrates rapidly.
This capacitance method is used to measure water dew point ranging from -150 to 70 0F with a response time of less than 5 seconds for a 63% step change in moisture content. As with the electrolysis method, contamination by pipe-scale, carbon, salt, and conductive liquids (glycol, methanol) can impede measurement. The sensor is not harmed by liquid slugs of condensate, methanol, glycol, and water. Proper cleaning restores the sensor.
4.2.4 Titration Method
The water content is determined by titration using a water specific reagent (usually Karl Fischer reagent). The sample gas enters the reaction cell, bubbles through a small known quantity of liquid reagent (0.5 mL), and exits via a reagent trap, a pressure reducing regulator, and finally a flow meter. A pair of platinum electrodes, sense the end point of the titration (when the entering water vapor has exhausted the batch of liquid reagent). Then a fresh batch of reagent is injected into the reaction cell by pump.
Electronic circuitry measures the time between end points and the sample gas flow-rate and pressure then it computes and then displays the water content. The entire cycle takes about two seconds.
Karl Fischer reagent is inert to hydrocarbons, carbon oxide, glycol, amines, and most sulfur compounds, e.g., odorizing mercaptans. Hydrogen sulfide will cause the moisture titrator to read high by 0.7 lb H2O/MMscf per grain H2S/100 scf.
Fig. 4-4. Aluminum oxide humidity sensor.
Fig. 4-5. Karl Fischer titrator for moisture content in gas.
4.2.5 Conductivity Cell
The Hygromat measuring cell consists of two stainless steel plates separated and electrically insulated from each other by a ceramic layer. The ceramic layer has eight holes which are partially filled with a hygroscopic salt-glycerol solution. Water is absorbed reversibly by the hygroscopic solution until equilibrium is reached with the surrounding natural gas. In turn the conductivity of the salt-glycerol solution increases as water is absorbed and decreases when water is desorbed.
4.2.6 Dew-Point Tubes
The dew-point tube uses a sampling pump as in figure 4-6. A water detector tube is placed in the pump and 100 mL of pipeline gas is pulled through the tube. The detector tube figure 4-7, is filled with magnesium perchlorate contained in a fine silica gel. Water vapor is absorbed by magnesium perchlorate to produce an alkaline reaction that changes the color of the Hammet’s indicator (crystal violet). The water content in lb/MMscf is read directly from the length of the stain in the detector tube. Overall accuracy is +/- 25%. Alcohols, glycols, and amines cause high readings. The tube range is usually 6-80 lb H2O/MMscf.
4.2.7 Comparison of Methods
- The Bureau of mines dew point method is respected as the one defined by an ASTM standard. Equipment cost is low, but the method is labor intensive in that a single measurements requires approximately 15 minutes. Accuracy varies with operator skill. Uncertainties of +/- 1 0F are attainable for dew points above 32 0F but increase to +/- 4 0F for dew points between -80 to -100 0F.
- The popularity of the electrolytic moisture analyzers is increasing significantly. Their accuracy compares favorably with the dew-point method. Advantages include light weight, probability, continuous readings, fast response times, and ready interfacing with alarms and other process monitors. Improved methods of cleaning and recoating the P2O5 film reduce the most frequent disadvantages of contamination and washout.
- The aluminum oxide sensor is relatively recent and exhibits accuracies and advantages similar to the electrolytic analyzer. It is especially suitable for every dry gases. However, response time is slow and removal of contaminants is more difficult.
- The titrator equipment is relatively expensive but not readily portable. The main advantages are accuracy (3% of reading) and immunity to contaminants, such as glycols and alcohols. The chief disadvantage is the hazardous nature of the Karl Fischer reagents, which creates a disposal problem.
- The advantages of the conductivity method include long-term stability. The 63% response time to sudden changes in gas humidity varies from 5 to 30 minutes depending on gas flow rate and pressure. One disadvantage is that conductivity varies with temperature and so the meter must be kept at constant temperature.
- Detector tubes provide inexpensive approximate estimate. They can be used by nontechnical personnel with minimum training.
Fig. 4-6. Sampling equipment and methods for dew-point tubes.
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Fig. 4-7. Dew-point tube readings.
4.3 Water Content of Natural Gases
The saturated water content of a gas depends on pressure, temperature, and composition. The effect of composition increases with pressure and is particularly important if the gas contains CO2 and/or H2S. For lean, sweet natural gases containing over 70% methane and small amounts of heavy hydrocarbons, generalized pressure-temperature correlations are suitable for many applications.
4.3.1 Water Content of Sweet Gases
Fig. 4-8 is an example of one such correlation which has been widely used for many years in the design of “sweet” natural gas dehydrators. The gas gravity correlation should never be used to account for the presence of H2S and CO2 and may not always be adequate for certain hydrocarbon effects, especially for the prediction of water content at pressures above 1500 psia. The hydrate formation line is approximate and should not be used to predict hydrate formation conditions.
Example 4-1 — Determine the saturated water content for a sweet lean hydrocarbon gas at 150°F and 1,000 psia.
From Fig. 4-8, W = 220 lb/MMscf.
For a 26 molecular weight gas, Cg = 0.98 (Correction for gas gravity- fig. 4-8)
W = (0.98)(220) = 216 lb/MMscf
For a gas in equilibrium with a 3% brine, Cs = 0.93 (Correction for salinity fig. 4-8)
W = (0.93)(220) = 205 lb/MMscf
Fig. 4-8. Water Content of sweet Hydrocarbon Gas. McKetta and Wehe, 1958; GPSA, 1987.
4.3.2 Water Content of High CO2/H2S Gases
Saturated water content of pure CO2 and H2S can be significantly higher than that of sweet natural gas, particularly at pressures above about 700 psia at ambient temperatures.
Corrections for H2S and CO2 should be applied when the gas mixture contains more than 5% H2S and/or CO2 at pressures above 700 psia. These corrections become increasingly significant at higher concentrations and higher pressures.
Below 40% acid gas components, one method of estimating the water content uses Eq 4-1 and Fig. 4-8, 4-9, and 4-10.
W = yHC WHC + yCO2 WCO2 + yH2SWH2S Eq 4-1
Where
W = Water content of gas, lb/MMscf
yHC = Mole fraction of hydrocarbon in the gas phase
WHC = Water content in hydrocarbon gas, lb/MMscf
YCO2 = Mole fraction of CO2 in the gas phase
WCO2 = Effective water content in CO2 gas, lb/MMscf
yH2S = Mole fraction of H2S in the gas phase
WH2S= Effective water content in H2S gas, lb/MMscf
Fig. 4-9. Effective Water Content of H2S in Natural Gas Mixtures vs. Temperature at Various Pressures
Fig. 4-10. Effective Water Content of CO2 in Natural Gas Mixtures vs. Temperature at Various Pressures
Note that Fig. 4-9 and 4-10 provide values for what is termed the “effective” water content of CO2 and H2S in natural gas mixtures for use only in Eq 4-1. These are not pure CO2 and H2S water contents.
A second method is Fig. 4-11. The CO2 is converted to equivalent H2S, using the factor 70%.
Example 4-2 — Determine the saturated water content of an 80% C1, 20% CO2 mixture at 160 °F and 2000 psia. The experimentally determined water content was 172 lb/MMscf.
Method One
WHC = 167 lb/MMscf (Fig. 4-8)
WCO2 = 240 lb/MMscf (Fig. 4-10)
W = (0.80 x 167) + (0.20 x 240)
=182 lb/MMscf
Method Two
First the composition must be converted for use with Fig. 4-11.
yH2S (pseudo) = 0.70 (yCO2 ) = 0.70 (0.20) = 0.14
Enter the left side of Fig. 4-11 at 160°F and move to the % H2S Equivalent line (14%). Proceed vertically upward to the Pressure, psia line (2000 psia), and move horizontally to the left to Water Content Ratio scale (ratio of 1.16).
W = (1.16)(167) = 194 lb/MMscf
Fig.4-11 . Calculated Water Content of Acid Gas Mixtures.
4.4 Prediction of Temperature drop due to pressure drop
Figure 4-12 can be used to get a quick approximate solution for the temperature drop of a natural gas stream (accuracy is +/- 5%.). For example, if the initial pressure is 4,000 psi and the final pressure is 1,000 psi, P is 3,000 psi, the change in temperature is 80°F. This curve is based on a liquid concentration of 20 bbl/MMscf. The greater the amount of liquid in the gas the lower the temperature drop, that is, the higher the calculated final temperature. For each increment of 10 bbl/MMscf there is a correction of 5°F. For example, if there is no liquid, the final temperature is 10°F cooler (the temperature drop is 10°F more) than indicated by Figure 4-12.
Example 4-3: Determine the temperature drop across a choke
Given: A well with a flowing tubing pressure of 4000 psi and 20 bbl of hydrocarbon condensate and a downstream back pressure of 1000 psi.
Solution: Initial pressure = 4000 psi
Final pressure = 1000 psi , P = 3000 psi
From Figure 4-12 correlation; intersect initial pressure = 4000 and ΔP-3000 read ΔT = 800F.
4.5 Hydrates in Natural Gas Systems
A hydrate is a physical combination of water and other small molecules to produce a solid which has an “ice-like” appearance but possesses a different structure than ice. Their formation in gas and/or NGL systems can plug pipelines, equipment, and instruments, restricting or interrupting flow.
There are three recognized crystalline structures for such hydrates. Where, water molecules build the lattice and hydrocarbons, nitrogen, CO2 and H2S occupy the cavities. Smaller molecules (CH4, C2H6, CO2, H2S) stabilize a body-centered cubic called Structure I. Larger molecules (C3H8, i-C4H10, n-C4H10) form a diamond-lattice called Structure II.
Normal paraffin molecules larger than n-C4H10 do not form Structure I and II hydrates as they are too large to stabilize the lattice. However, some isoparaffins and cycloalkanes larger than pentane are known to form Structure H hydrates.
Gas composition determines structure type. Mixed gases will typically form Structure II. From a practical viewpoint, the structure type does not affect the appearance, properties, or problems caused by the hydrate. It does however, have a significant effect on the pressure and temperature at which hydrates form. Structure II hydrates are more stable than Structure I. This is why gases containing C3H8 and i-C4H10 will form hydrates at higher temperatures than similar gas mixtures which do not contain these components. The effect of C3H8 and i-C4H10 on hydrate formation conditions can be seen in Fig. 4-15. At 1000 psia, a 0.6 sp. gr. gas (composition is shown in Fig. 4-15) has a hydrate formation temperature which is 12-13°F higher than pure methane.
Fig.4-13. Hydrate structures, typical hydrate plugging, and illustration of hydrate formation.
Fig. 4-14. Conditions for Hydrate Formation for Light Gases
Fig. 4-15. Pressure-Temperature Curves for Predicting Hydrate Formation. Katz.1945.GPSA 1987.
The presence of H2S in natural gas mixtures results in a substantially warmer hydrate formation temperature at a given pressure. CO2, in general, has a much smaller impact and often reduces the hydrate formation temperature at fixed pressure for a hydrocarbon gas mixture.
4.5.1 Conditions which affect hydrate formation are:
Primary Considerations
• Gas or liquid must be at or below its water dew point or saturation condition (Note: liquid water does not have to be present for hydrates to form)
• Temperature
• Pressure
• Composition
Secondary Considerations
• Mixing
• Kinetics
• Physical site for crystal formation and agglomeration such as a pipe elbow, orifice, thermowell, or line scale
• Salinity
In general, hydrate formation will occur as pressure increases and/or temperature decreases to the hydrate formation condition.
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4.5.2 Prediction of Sweet Natural Gas Hydrate Conditions
Fig. 4-14, based on experimental data, presents the hydrate pressure-temperature equilibrium curves for pure methane, ethane, propane, and for a nominal 70% ethane 30% propane mix.
Fig. 4-15 through 4-19, based on gas gravity, may be used for first approximations of hydrate formation conditions and for estimating permissible expansion of sweet natural gases without the formation of hydrates.
The conditions at which hydrates can form are strongly affected by gas composition.
Example 4-4- Find the pressure at which hydrate forms for a gas with the following composition. T = 50°F.
Table 4-1. Gas composition and molecular weight calculation for Example 4-4.
Gas specific gravity = MWgas/MWair = 20.08/28.964 = 0.693
From Fig. 4-15 at 50°F.
P = 320 psia for 0.7 gravity gas
Example 4-5— The gas in Example 4-4 is to be expanded from 1,500 psia to 500 psia. What is the minimum initial temperature that will permit the expansion without hydrate formation?
The 1,500 psia initial pressure line and the 500 psia final pressure line intersect just above the 110°F curve on Fig. 4-17. Approximately 112°F is the minimum initial temperature.
Example 4-6 — How far may a 0.6 gravity gas at 2,000 psia and 100°F be expanded without hydrate formation?
On Fig. 4-16 find the intersection of 2,000 initial pressure line with the 100°F initial temperature curve. Read on the x-axis the permissible final pressure of 1100 psia.
Fig. 4- 16. Permissible Expansion of a 0.6-Gravity Natural Gas Without Hydrate Formation
Example 4-7 — How far may a 0.6 gravity gas at 2,000 psia and 140°F be expanded without hydrate formation?
On Fig. 4-16, the 140°F initial temperature curve does not intersect the 2,000 psia initial pressure line. Therefore, the gas may be expanded to atmospheric pressure without hydrate formation.
Conditions predicted by Fig. 4-15 through 4-19 may be significantly in error for compositions other than those used to derive the charts. For more accurate determination of hydrate formation conditions, the following procedures should be followed. In addition, fig. 4-15 through 4-19, do not account for liquid water and liquid hydrocarbons present or formed during the expansion. These can have a significant effect on the outlet temperature from the pressure reduction device.
Fig. 4- 19. Permissible Expansion of a 0.9-Gravity Natural Gas Without Hydrate Formation
4.5.3 Hydrate Prediction Based on Composition for Sweet Gases
1- Katz’s Graph
The first method is for approximate and fast prediction, using of Katz graph represented in figure 4-15. On the other hand, for below 1,000 psi (70 bar), the figure can be approximated by:
t(°F) = −16.5 – [6.83/(SpGr)2] + 13.8 ln[P(psia)] Eq 4-9.
2- Katz’s Vapor solid equilibrium method
Several correlations have proven useful for predicting hydrate formation of sweet gases and gases containing minimal amounts of CO2 and/or H2S. The most reliable ones require a gas analysis. The Katz method utilizes vapor solid equilibrium constants defined by the Eq 4-10.
Kvs = y / xs Eq 4-10
Where
Kvs = vapor/solid equilibrium K-value
y = mole fraction in the gas phase
xs = mole fraction in the solid phase
WARNING: Not good for pure components – only mixtures.
The Katz’s correlation is not recommended above 1000-1500 psia, depending on composition.
The applicable K-value correlations for the hydrate forming molecules (methane, ethane, propane, isobutane, normal butane, carbon dioxide, and hydrogen sulfide) are shown in Fig. 4-20 to 4-26. Normal butane cannot form a hydrate by itself but can contribute to hydrate formation in a mixture.
For calculation purposes, all molecules too large to form hydrates have a K-value of infinity. These include all normal paraffin hydrocarbon molecules larger than normal butane.
Nitrogen is assumed to be a non-hydrate former and is also assigned a K-value of infinity.
The Kvs values are used in a “dewpoint” equation to determine the hydrate temperature or pressure. The calculation is iterative and convergence is achieved when the following objective function (Eq 4-3) is satisfied.
∑y/Kvs = 1.0 Eq. 4-11
Fig. 4- 20. Kvs. Vapor solid equilibrium constants for Methane.
Fig. 4- 21. Kvs. Vapor solid equilibrium constants for Ethane.
Fig. 4- 22. Kvs. Vapor solid equilibrium constants for Propane.
Fig. 4- 23. Kvs. Vapor solid equilibrium constants for Iso-Butane.
Fig. 4- 24. Kvs. Vapor solid equilibrium constants for N-Butane.
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Fig. 4- 25. Kvs. Vapor solid equilibrium constants for carbon dioxide.
Fig. 4- 26. Kvs. Vapor solid equilibrium constants for hydrogen sulfide.
Example 4-8 — Calculate the pressure for hydrate formation at 50°F for a gas with the following composition.
Table. 4-2. Example. 4-8.
The ∑y/Kvs value is slightly over than 300 psia, by iterpolating linearly, = 1.0 @ 305 psia.
Hydrate pressure at temperature 50 0F = 305 psia
Third and fifth columns are values obtained from chart for each component at temperature and pressure values.
Fourth column contains the results of dividing Mole fraction of gas by third column.
Sixth column contains the results of dividing Mole fraction of gas by fifth column.
3- Motiee (1991) suggested the following equation for hydrate temperature prediction:
T(0F) = -238.24469 + 78.99667 log P(psi) – 5.352544 (log P(psi) )2 +
349.473877 y – 150.854675 y2 – 27.604065 y log P(psi) Eq. 4-12
Where ᵞ is gas specific gravity.
This equation is well known and widely used in the oil and gas industry because of its accuracy for natural gas mixtures.
4- Recently (2009), a new correlation developed by Bahadori and Vuthaluru, with specific gravities from 0.55 to 1, shows the best efficiency. The equation is suitable for estimating the HFT, especially for natural gas mixtures:
T(k) = AyB (ln P(KPa))C Eq. 4-13
where:
A=194.681789
B=0.044232
C=0.189829
0F = (0K - 273.15) x 9/5 + 32
0K = (0F - 32) x 5/9 + 273.15
(0C=0K-273.15)
psi = 6.895 KPa,
4.5.4 Hydrate Predictions for High CO2/H2S Content Gases
The Katz method of predicting hydrate formation temperature gives reasonable results for sweet paraffin hydrocarbon gases. The Katz method should not be used for gases containing significant quantities of CO2 and/or H2S despite the fact that Kvs values are available for these components. Hydrate formation conditions for high CO2/H2S gases can vary significantly from those composed only of hydrocarbons. The addition of H2S to a sweet natural gas mixture will generally increase the hydrate formation temperature at a fixed pressure. A method by Baille & Wichert for predicting the temperature of high H2S content gases is shown in Fig. 4-27.
Fig.4-27. Hydrate Chart for Gases Containing H2S
4.6 Hydrate Prevention
Hydrate prevention is accomplished by keeping the:
1- Operating conditions must remain out of the hydrate-formation zone by heating or temperature control.
2- Hydrate point must be maintained below the operating conditions of the system by chemical treatment.
Two common methods of hydrate-formation prevention are:
1- Temperature control
2- Chemical injection
Example 4-9 — Estimate the hydrate formation temperature at 610 psia of a gas with the following analysis using Fig. 4-27.
Table. 4-3. Solution of example 4-9.
MW = 19.75
Sp.Gr (ᵧ) = 0.682
Solution Steps:
1. Enter left side of Fig. 4-27 at 600 psia and proceed to the H2S concentration line (4.18 mol%)
2. Proceed down vertically to the specific gravity of the gas (ᵞ = 0.682)
3. Follow the diagonal guide line to the temperature at the bottom of the graph (T = 63.5°F).
4. Apply the C3 correction using the insert at the upper left.
Enter the left hand side at the H2S concentration and proceed to the C3 concentration line (0.67%). Proceed down vertically to the system pressure and read the correction on the left hand scale (–2.7°F)
Note: The C3 temperature correction is negative when on the left hand side of the graph and positive on the right hand side.
TH = 63.5 −2.7 = 60.8°F
Fig. 4-27 was developed based on calculated hydrate conditions using the Peng-Robinson EOS. It has proven quite accurate when compared to the limited amount of experimental Mole Fraction data available. It should only be extrapolated beyond the experimental data base with caution.
4.6.1 Adding Heat
Adding heat is effective because hydrates normally do not occur above 70 0F.
It offers a simple and economical solution for land and offshore facilities (if waste heat is available).
Flow stream is preheated, either through an indirect line heater or heat exchanger, before passing through a choke. Flow stream is then reheated to maintain the temperature above the hydrate formation temperature.
A major drawback in offshore installations is that it is almost impossible to maintain flowline temperatures significantly above the water temperature if the flowlines extend more than a few hundred feet under water. Thus, either the “free water” must be separated while still at temperature or an alternate method
selected.
4.6.1.1 Temperature Control
Wellhead Indirect Heaters
An indirect heater is used to heat gas to maintain temperatures above that of the hydrate formation.
It consists of an atmospheric vessel containing a fire tube (usually fired by gas, steam, or heating oil) and a coil (designed to withstand shut in tubing pressure “SITP”) that is heated by the intermediate fluid (usually water) and the fluid is heated. The fire tube and coil are immersed in a heat transfer fluid (normally water), and heat is transferred to the fluid in the coil.
Figure 4-28 shows a typical heater installation at the wellhead.
Long-Nose Heater Choke (Figure 4-29).
A long body choke installed in the indirect heater to position the choke orifice within the indirect heater bath. Since the walls of the choke orifice are heated by the water bath, hydrates will not form in the orifice and cause plugging.
Flowline indirect Heaters
Flowline heaters differ from wellhead heaters in purpose only.
The purpose of a wellhead heater is to heat the flow stream at or near the wellhead where choking or pressure reduction occurs.
The purpose of a flowline heater is to provide additional heat if required.
The design is the same as an indirect heater except that the choke, shut-in, and relief equipment are seldom used.
System Optimization
System operation has to be optimized before heaters can be effectively designed and located.
Heat requirements that appear to be large can often be reduced to minimal values or even eliminated by revising the mode of operation. For example: Fields having multiple producing wells can be combined to use higher flowing temperatures thus minimizing heater requirements.
If reducing the gas stream pressure is necessary, it is generally more efficient to do so at a central point where the necessary heater fuel gas can be obtained from separators or scrubbers.
Requires flowline wall thickness to be increased so as to withstand wellhead SITP.
An alternative is to install wellhead shut-down valves and flowline high pressure switches.
Downhole Regulators
Downhole regulators are feasible for high capacity gas wells at locations where certain risks to other downhole equipment are acceptable. The theory behind the use of a downhole regulator is that the pressure drop from flowing pressure to near-sales line pressure is taken downhole where the formation temperature is sufficient to prevent hydrate formation. The tubing string above the regulator then acts as a subsurface heater. Calculations involved in downhole regulator design are rather involved. They depend on characteristics such as:
Wellbore configuration, flowing downhole pressures and temperature, and well depth.
Although shortcut procedures are available to estimate the feasibility of downhole regulators, tool company representatives can provide detailed design information.
4.6.2 Chemical Injection
The formation of hydrates can also be prevented by dehydrating the gas or liquid to eliminate the formation of a condensed water (liquid or solid) phase. In some cases, however, dehydration may not be practical or economically feasible. In these cases, chemical inhibition can be an effective method of preventing hydrate formation. Chemical inhibition utilizes injection of thermodynamic inhibitors or low dosage hydrate inhibitors (LDHIs). Thermodynamic inhibitors are the traditional inhibitors (i.e., one of the glycols or methanol), which lower the temperature of hydrate formation. LDHIs are either kinetic hydrate inhibitors (KHIs) or antiagglomerants (AAs).
They do not lower the temperature of hydrate formation, but do diminish its effect. KHIs lower the rate of hydrate formation, which inhibits its development for a defined duration.
AAs allow the formation of hydrate crystals but restrict them to sub-millimeter size.
Fig. 4-28. Wellhead indirect heater schematic.
4.6.2.1 Thermodynamic Inhibitors
Inhibition utilizes injection of one of the glycols or methanol into a process stream where it can combine with the condensed aqueous phase to lower the hydrate formation temperature at a given pressure.
Both glycol and methanol can be recovered with the aqueous phase, regenerated and re-injected. For continuous injection in services down to –40°F, one of the glycols usually offers an economic advantage versus methanol recovered by distillation.
At cryogenic conditions (below –40°F) methanol usually is preferred because glycol’s viscosity makes effective separation difficult.
Ethylene glycol (EG), diethylene glycol (DEG), and triethylene glycol (TEG), glycols have been used for hydrate inhibition. The most popular has been ethylene glycol because of its lower cost, lower viscosity, and lower solubility in liquid hydrocarbons.
Freezing point of aqueous methanol solution is given in fig. 4-30.
Hydrate inhibitors are used to lower the hydrate formation temperature of the gas.
Recovery and regeneration steps are used in all continuous glycol injection projects and in several large-capacity methanol injection units.
Injection of hydrate inhibitors should be considered for the following applications:
• Pipeline systems in which hydrate trouble is of short duration
• Gas pipelines that operate at a few degrees below the hydrate formation temperature
• Gas-gathering systems in pressure-declining fields
• Gas lines in which hydrates form as localized points
Methanol and the lower molecular weight glycols have the most desirable characteristics for use as hydrate inhibitors.
When hydrate inhibitors are injected in gas flowlines or gathering systems, installation of a free-water knockout (FWKO) at the wellhead proves to be economical in nearly every case.
Removing the free water from the gas steam reduces the amount of inhibitor required.
To be effective, the inhibitor must be present at the very point where the wet gas is cooled to its hydrate temperature. For example, in refrigeration plants glycol inhibitors are typically sprayed on the tube-sheet faces of the gas exchangers so that it can flow with the gas through the tubes. As water condenses, the inhibitor is present to mix with the water and prevent hydrates. Injection must be in a manner to allow good distribution to every tube or plate pass in chillers and heat exchangers operating below the gas hydrate temperature.
Table 4-4 lists some physical properties of methanol and the lower molecular weight glycols.
The inhibitor and condensed water mixture is separated from the gas stream along with a separate liquid hydrocarbon stream. At this point, the water dewpoint of the gas stream is essentially equal to the separation temperature. Glycol-water solutions and liquid hydrocarbons can emulsify when agitated or when expanded from a high pressure to a lower pressure, e.g., JT expansion valve. Careful separator design will allow nearly complete recovery of the diluted glycol for regeneration and reinjection.
Fig. 4-29. Indirect heater and long-Nose choke.
Fig. 4-30. Freezing Points of Aqueous Methanol Solutions
Methanol Injection Considerations
Methanol is well-suited for use as a hydrate inhibitor because it is:
Noncorrosive
Nonreactive chemically with any constituent of the gas
Soluble in all proportions in water
Volatile under pipeline conditions
Reasonable in cost
Of a vapor pressure greater than that of water
Methanol is injected by means of an injection pump (3 in Figure 4-31) into the flowline upstream of the choke or pressure control valve (2). A temperature controller may be installed to measure the temperature of the gas in the low-pressure flowlineand adjusts the methanol rate accordingly.
Fig. 4-31. Methanol injection system at wellhead.
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Fundamentals of Oil and Gas Processing
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Basics of Gas Field Processing
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Chapter 4 - Part 2
Glycol Injection Considerations
Glycol has a relatively low vapor pressure and thus does not evaporate into the vapor phase as readily as methanol. The solubility of glycol in liquid hydrocarbons is relatively low. For the above reasons, glycol can be more economically recovered, thus reducing the operating expenses below those of methanol systems. The injection part of the system (items 1 to 5 in Figure 4.32) is similar to the methanol injection system. Taking into consideration that, the viscosities of ethylene glycol and its aqueous solutions increase significantly as temperature decreases. Additional equipment in the glycol system is for recovering and reclaiming the glycol. A three-phase separator (6) separates the water and glycol from the hydrocarbon phases. The water–glycol solution in the separator is sent to the reboiler (7) while gas is delivered to the sales line, and the hydrocarbon condensate is dumped to the condensate storage tanks.
In the reboiler, excess water is boiled away from the glycol. The glycol reconcentrated in the reboiler is then available again for injection into the gas stream. Separation of the glycol water phase from the hydrocarbon-liquid field requires a temperature above 700F and a residence time of 10 to 15 minutes.
Nozzle Design (Figure 4-33)
Due to the vapor pressure of glycol, a fine, well-distributed mist is required to obtain adequate mixing with the gas to ensure optimum results, thus spray nozzles are normally used. Nozzle selection is a major consideration in the design of cold separation facilities or plants using glycol injection. Glycol injection normally takes place just upstream of a heat exchanger or chiller where gas is being chilled. Proper nozzle selection will ensure that the glycol spray covers the tube sheet. 100 to 150 psi differential pressure at the nozzle is sufficient to atomize the glycol.
Process stream velocities should be at least 12 ft/s.
Fig. 4-32. Glycol injection system at wellhead.
Fig. 4-33. Schematic of a spray nozzle used in glycol injection.
Glycol Selection General Guideline
The three glycols normally used to prevent the formation of hydrates are:
• Ethylene glycol (EG)
• Diethylene glycol (DEG)
• Triethylene glycol (TEG)
Selection of a glycol depends on the composition of the hydrocarbon flow stream as follows:
1- If glycol is to be injected into a natural gas transmission line where glycol recovery is of less importance than hydrate protection, ethylene glycol is the best choice because it produces the greatest hydrate depression and has the highest vapor pressure of any of the glycols.
2- If glycol is to be injected into a unit where it will contact hydrocarbon liquids, ethylene glycol is preferred because it has the lowest solubility in high molecular weight hydrocarbons.
3- If vaporization losses are severe, either diethylene or triethylene glycol are the best choice because both have a lower vapor pressure. Sometimes diethylene glycol is used if there is a combined loss of both gas vaporization and liquid solubility.
4- The freezing point of the glycol solution must be lower than the lowest temperature expected in the system. In inhibitor service, glycol concentrations are usually maintained at 70 to 75 weight percent because freezing of the glycol is not a problem at this concentration.
5- Reboiler temperature is dependent on the type of glycol and its concentration.
6- Temperature should be maintained at a level equal to the boiling point of the desired solution.
7- Boiling points for the three glycol types are plotted in Figures 4-34, and 4-35.
For example, from Figure 4-34 the reboiler temperature should be set at 240 0F in order to produce about 70 weight percent ethylene glycol solution at atmospheric pressure (N.B. 760 mm, Hg= 1 atm. Pressure.). Thermal degradation can occur if the boiling point of the pure glycol is exceeded; it should therefore be avoided. Glycol losses for the two-phase gas condensate systems are normally estimated at 1 to 2 gallons per 100 barrels of hydrocarbon liquid produced. Vaporization into the gas stream and solution into the hydrocarbon liquid usually cause only a small portion of the total loss.
The most significant causes of glycol losses are leakage and carryover with the hydrocarbon liquid. Loses also occur from vaporization and carry over in the reboiler.
Fig. 4-34. Boiling and condensation points of Ethylene and Di-ethylene glycol.
Fig. 4-35. Boiling and condensation points of Tri-ethylene glycol.
Chemical Injection System
The three parts of an injection system are:
Pump, meter, and control system
Single-Point Chemical Injection
A single pump, meter, and control system service one injection point.
Disadvantages are:
1- Limited turn-down capability and increased life-cycle cost
2- Weight and space increase as injection points increase
Multi-Point Chemical Injection
A shared pump and multiple meter and control devices servicing multiple injection points.
Advantages:
1- Increased turn-down capacity.
2- Per well capital investment decreases as the number of wells increases
3- Injection points are easily added
4- Lower weight and space requirements for higher quantity well applications
Disadvantages
1- Instrumentation intensive, and multiple control loops required
2- Requires variable speed for fixed crank pumps
3- Experiences high-pressure drops from header to recycle line
Metering Pump Types:
Diaphragm Pumps
Advantages
Hermetically sealed, no contamination to atmosphere
Long-life diaphragms typically greater than 2 years continuous duty (20,000 hours)
Long-life of hydraulic plunger seals, typically greater than 2 years continuous duty.
Internal hydraulic relief
Maximum safeguard to environment and personnel safety.
Disadvantages
Higher purchase price
More complex maintenance required
Plunger Pumps
Advantages
Lower purchase price
Less complicated maintenance (easier to understand)
Disadvantages
Plunger packing service life typically less than 2000 hours
Friction between plunger and packing
Comparison of Hydrate Prevention Methods
Overview
The four methods (indirect heaters, methanol injection, glycol injection, and downhole regulators) discussed above are proven safe and reliable.
Evaluation should consider:
1- Development of CAPEX and OPEX (including chemicals and fuel)
2- Space needs (especially in offshore operations) and operating hazards.
Heaters
1- Capital costs and the fuel expense of heaters are relatively large, and it is difficult to maintain a clean, reliable fuel supply to remote heater locations.
2- Indirect heater requires a large amount of space.
3- Fire boxes with proper flame arrestors have minimized the hazards from fired equipment, but they should be bought with strict attention paid to detailed design.
Chemical Injection
Advantages and disadvantages of methanol injection and glycol injection are listed in Table 4-5.
The use of methanol requires only a free-water separator and a suitable means for injection and atomizer, whereas the use of glycol requires a free-water separator plus a gas–liquid separator and a glycol reconcentration unit at the point of recovery downstream.
Downhole Regulators
No routine service is required on downhole regulators, but a wireline service company must be used each time the pressure drop has to be changed and when the regulator is removed. A well with a downhole regulator may require injection of methanol or glycol when it is brought back online after a shut-in until the flow and temperature stabilize. After a well declines to less than allowable production, the downhole regulator will have to be removed, and another form of hydrate prevention may prove necessary.
Downhole regulators do not present special safety hazards, but because work with regulators involves working in the well, losing the well is always a danger.
Inhibitor Advantages Disadvantages
Methanol Relatively low initial cost Minimal equipment
Simple system
High operating cost
Hauling to site necessary
Disposal procedures and precautions for water-methanol mixture
Glycol Usually lower operating cost than methanol when both systems recover chemical
Simple system High initial cost
Hauling to site necessary
Large loss if line breaks
Table 4-5 Methanol and Glycol Injection Comparisons
Summary of Hydrate Prevention Methods:
The methanol injection system is often used for temporary hydrate prevention service in small installations.
Larger installations are favored for indirect heaters or glycol injection systems.
Downhole regulators are most useful in large high-pressure reservoirs in which excess pressure is available and the reservoir pressure is not expected to decline rapidly.
Table 4-6 contains a summary comparison of the above methods.
4.6.3 Hydrate Inhibition with Methanol and Glycols
The amount of chemical inhibitor required to treat the water in order to lower the hydrate formation temperature may be calculated from the Hammerschmidt equation:
ΔT = KWR / M (100-WR) Eq. 4- 14
Where
ΔT is the depression in hydrate formation temperature (0F),
WR is weight percent of inhibitor for water treatment,
K is a constant that depends on the type of inhibitor (table 4-7), and
M is the molecular weight of the inhibitor (table 4-7)
Table. 4-7. M and K values for hydrate inhibitors.
Once the required inhibitor concentration has been calculated, the mass of inhibitor required in the water phase may be calculated from the following equation:
Amount of inhibitor (mI) for water phase for each lb/MMscf
mI = WR X lbH2O/MMscf /(1- WR) Eq. 4-15
Or in general equation
mI = WR X mH2O / (WL − WR) Eq 4-16
mI = Amount of Inhibitor (lb)
mH2O = Amount of water (lb)
WR = weight percent of inhibitor for water treatment,
WL = Lean inhibitor concentration (for example, 80% EG; WL =0.8, 100% methanol; WL=1, etc.)
Fig. 4-36 . Ratio of methanol vapor composition to methanol liquid composition
Determination of Total Inhibitor Required
Total inhibitor required =
Inhibitor required for free water + Inhibitor lost to vapor phase + Inhibitor soluble in condensate
Where inhibitor lost to vapor phase is determined from Figure 4-36. (Methanol lost to the vapor phase), while glycol vaporization losses are generally very small and are typically ignored in calculations.
Inhibitor soluble in the condensate is determined from Figure 4-37 and a value of approximately 0.5% can be used. Solubility of EG in the liquid hydrocarbon phase is extremely small. A solubility of 0.3 lb per 1000 gal. (U.S.) of NGL is often used for design purposes.
Engineering Data Book “GPSA 2004”, recommended that Eq 4-15, should not be used beyond 20-25 wt% for methanol and 60-70 wt% for the glycols. For methanol concentrations up to about 50%, the Nielsen-Bucklin equation (4-17) provides better accuracy:
ΔT 0F =−129.6 ln (zH2O) Eq 4-17
where
z = mole fraction in the liquid phase
zH2O is mole fraction of water
Mole fraction of inhibitor, zI = 1 - zH2O.
WR = (MI X zI ) / (MI X zI + 18 X zH2O) Eq.4-18
MI = Molecular weight of inhibitor
For methanol, WR = (32 X zI )/ (32 X zI + 18 X zH2O). (Examples 4-10, and 4-11 will explain).
Fig. 4-37. Solubility of Methanol in Paraffinic Hydrocarbons vs. Temperature at Various Concentrations
Example 4-10: Determining the Amount of Methanol Required in a Wet Gas Stream
Given:
Flow rate
Gas= 20 MMscfd (Sp.Gr = 0.600)
Condensate = 800 bpd (60 0API/ SP.Gr.=0.739) = 40 bbl/MMscf
Produced Water = 60 Bwpd (Sp.Gr.= 1.03) = 3 bbl/MMscf
Upstream choke
FWHP = 3000 psia (P1)
FWHT = 100 0F (T1)
Down stream choke (cold line)
FWHP = 2000 psia (P2)
FWHT = 60 0F (T2)
Determine: Calculate the total methanol required to prevent hydrates from forming.
Fig. 4-38. Example 4-9.
Solution:
1. The amount of water that will be condensed is determined from McKetta-Wehe (Figure 4-8), assuming the gas is saturated at wellhead conditions.
Water content upstream choke @ 3,000 psia & 100 0F = 32.0 Ib/MMscf
Water content downstream choke @ 2,000 psia & 60 0F = 11.5 lb/MMscf
Water Condensed =(32.0 – 11.5) = 20.5 lb/MMscf
Produced Water = (1.03 X 2.205 X 159) = (361 lb/bbl) X 3 = 1083 lb/MMscf
Total water = (20.5 + 1083) =1103.6 Ib/MMscf
2. From pressure-temperature curve, the hydrate formation temperature is 68 0F (refer to Figure 4-15), or according to equation 4-12, or Equation 4-13 Hydrate formation temperature is 66 0F.
(In this calculation, we will consider the value 68 0F)
The required dewpoint depression then is 68 0F– 60 0F = 8 0F
3. The concentration of methanol required in the liquid water phase from Equation 4-14 is:
8 0F = 2335 WR / 32 (100 – WR)
Rearranging and solving for WR = 9.892% = 0.09892
4. Therefore, from equation 4-15, the estimated methanol required in the liquid water phase is:
mI = 0.09892 x 1103.65 / (1 – 0.09892 ) = 121 lb/MMscf
or we can use equation mI = WR X mH2O / (WL − WR) Eq 4-16
5. From Figure 4-36, the methanol that will flash into the vapor phase at 2000 psia and 60 0F is
= n lbs.methanol/MMscf / WT% methanol in water phase = 1.52
6. Therefore, the methanol in the vapor phase (n) is:
n= 1.52 X 9.892% =15 lb/MMscf
7. A barrel of our condensate weighs:
= [0.739 x 2.205 (lb per liter water) x 159 (liter per barrel)] = 259 lb/bbl
8. Therefore, the approximate amount of methanol soluble in the condensate or liquid hydrocarbon phase (assuming a 0.5% solubility by weight) = 0.005 x 259 x 40 = 52 lb/MMscf
9. Thus, the total amount of methanol required is:
Total = liquid water phase (121) + Vapor phase (15) + Soluble in condensate (52) =188 lb/MMscf
Total = (188 lb/MMscf) x (20 MMscfd) = 3760 lb/day
Note that for gas-condensate wells producing a reasonable or high amount of condensate, the amount of methanol soluble in the condensate is crucial to determining the amount needed.
Approximately 188 lb of methanol must be added so that approximately 121 lb will be dissolved into the water phase.
Since the specific gravity of methanol is 0.791 (at 68 0F), this is equivalent to:
= 188/ [0.791 x 2.205 (lb per liter water) ] = 108 liter/MMscf
= 28.5 gal/ MMscf
For 20 MMscfd, methanol injection will be 570 GPD (13.6 bbl/d).
Example 4-11—100 MMscf/d of natural gas leaves an offshore platform at 100°F and 1200 psia. The gas comes onshore at 40°F and 900 psia. The hydrate temperature of the gas is 65°F. Associated condensate production is 10 bbl/MMscf. The condensate has an API gravity of 50 and a MW of 140. Calculate the amount of methanol and 80 wt% EG inhibitor required to prevent hydrate formation in the pipeline.
Solution Steps:
Methanol
1. Calculate the amount of water condensed per day from McKetta-Wehe (Figure 4-8),
Win =53.0 lb /MMscf
Wout = 9.5 lb/MMscf
ΔW = 43.5 lb/MMscf
Water condensed = (100)( 43.5) = 4350 lb/day
The required dewpoint depression is 65 0F– 40 0F = 25 0F
2. Calculate required methanol inhibitor concentration from Eq 4-14 and 4-17.
ΔT = KWR / M (100-WR) Eq. 4- 14
25 0F = 2335 WR / 32 (100 – WR)
WR = 25% = 0.25
Using equation 4-17
ΔT 0F =−129.6 ln (zH2O) Eq 4-17
zH2O = 0.825 , i.e. zI = (1- 0.725 )= 0.175
From WR = (MI X zI ) / (MI X zI + 18 X zH2O) Eq.4-18
WR = (0.175 X 32) / (0.175 X 32 + 0.825 X 18) = 0.275
WR = 0.275, (use this value in subsequent calculations)
3. Calculate mass rate of inhibitor in water phase
Amount of inhibitor for water phase for each lb/MMscf
mI (lb)= WR X lbH2O/MMscf /(1- WR) Eq. 4-15
= 0.275 X 43.5 /(1-0.275) = 16.5 lb/MMscf = 1650 lb/day
4. Estimate vaporization losses from Fig. 4-36. @ 40°F and 900 psia, losses = 1.05 lb/MMscf
wt% MeOH
daily losses = (1.05)(100)(27.5) = 2890 lb/day
5. Estimate losses to hydrocarbon liquid phase from Fig.4-37. @ 40°F and 27.5 wt% MeOH, approximately = 0.2 mol% lb • mols of condensate per day– “note mole% not weight %”
6. A barrel of our condensate weighs: (API =50, i.e., Sp.Gr = 0.7796), molecular weight =140.
lb.Mole per day = 100 MMscfd X 10 bbl/MMscf X [0.7796 x 2.205 (lb per liter water) x 159 (liter per barrel)] /140 = = 1950 lb−mols/day.
7. lb • mols methanol = (1950)(0.002) = 3.9 lb • mols/day
lb methanol = (3.9)(32) = 125 lb/day
Total methanol injection rate = 1650 + 2890 + 125 = 4665 lb/day
Since the specific gravity of methanol is 0.791 (at 68 0F), this is equivalent to:
4665 / [0.791 x 2.205 (lb per liter water) ] = 2674 liter/day
= 26.74 Liter/ MMscf
Methanol left in the gas phase can be recovered by condensation with the remaining water in downstream chilling processes. Likewise, the methanol in the condensate phase can be recovered by water by downstream water washing.
80 wt% EG
1. Calculate required inhibitor concentration from Eq 4-14.
ΔT = KWR / M (100-WR) Eq. 4- 14
25 0F = 2335 WR / 62 (100 – WR)
W = 40% = 0.4
2. Calculate mass rate of inhibitor in water phase from Eq.4-16.
mI = WR X mH2O / (WL − WR) Eq 4-16
mI =0.4 X 4350 / (0.8 -0.4) = 4350 lb/day
Vaporization and liquid hydrocarbon losses are negligible.
Inhibitor losses represent a significant operating cost and can cause problems in downstream process units. Efficient inhibitor separation should be provided.
4.6.4 Low Dosage Hydrate Inhibitors (LDHIs)
LDHIs can provide significant benefits compared to thermodynamic inhibitors including:
1- Significantly lower inhibitor concentrations and therefore dosage rates. Concentrations range from 0.1 to 1.0 weight percent polymer in the free water phase, whereas alcohols can be as high as 50%
2- Lower inhibitor loss caused by evaporation, particularly compared to methanol
3- Reduced capital expenses through decreased chemical storage and injection rate requirements; and no need for regeneration because the chemicals are not currently recovered.
4- These are especially appropriate for offshore where weight and space are critical to costs
5- Reduced operating expenses in many cases through decreased chemical consumption and delivery frequency
6- Increased production rates, where inhibitor injection capacity or flowline capacity is limited
7- Lower toxicity
4.6.4.1 Kinetic Hydrate Inhibitors
KHIs were designed to inhibit hydrate formation in flowlines, pipelines, and downhole equipment operating within hydrate-forming conditions such as subsea and cold-weather environments. Their unique chemical structure significantly reduces the rate of nucleation and hydrate growth during conditions thermodynamically favorable for hydrate formation, without altering the thermodynamic hydrate formation conditions (i.e., temperature and pressure). This mechanism differs from methanol or glycol, which depress the thermodynamic hydrate formation temperature so that a flowline operates outside hydrate-forming conditions.
KHIs Compared to Methanol or Glycols
KHIs inhibit hydrate formation at a concentration range of 0.1–1.0 weight percent polymer in the free water phase. At the maximum recommended dosage, the current inhibition capabilities are 28°F of subcooling in a gas system and 20°F in an oil system with efforts continuing to expand the region of effectiveness.
For relative comparison, methanol or glycol typically may be required at concentrations ranging 20 to 50 weight percent respectively in the water phase.
KHI Screening Considerations
Although KHIs are applicable under most producing conditions, certain conditions must be considered when evaluating a potential application, which include water salinity, freezing conditions, hold time (i.e., period of effectiveness), water saturation, and high temperature processes.
• At water salinity levels greater than approximately 17% NaCL, the polymer may come out of solution, thereby reducing KHI effectiveness.
• A solution of KHI in water does not provide protection from freezing or icing conditions, neither in the line being treated nor in the KHI storage tank. If ambient temperatures are expected to fall below freezing, the KHI storage volume must be freeze-protected through the use of insulation on the container and piping or addition of antifreeze (typically ethylene glycol) to the KHI solution.
• A solution of KHI cannot be used for melting ice or hydrate plugs. It is recommended to have other strategies, such as a small quantity of ethylene glycol or methanol for remediation purposes in the event of a blockage.
• The KHI delivery system must be capable of providing sufficient dosage to achieve a hold time greater than the water residence time in the piping. Factors to consider include:
— The design basis duration of hydrate forming conditions for an unplanned shut-in.
— The potential for water to pool in low sections of piping (e.g., turn-down hydraulics, flowline profile, pigging frequency, flowline interconnects that are not used continuously) and dead legs.
— The seasonal duration of the cold point temperature below hydrate temperature, if applicable.
• If the gas is undersaturated with respect to water, the water in the KHI solution will evaporate and leave a high viscosity fluid. This can be addressed by using a more dilute KHI solution, or by changing the KHI carrier fluid to ethylene glycol.
• The KHI and water from the KHI solution will form separate phases if the inhibited fluid is above the lower critical solution temperature (LCST) of the KHI solution.
• The KHI polymer suffers degradation effects at temperatures above 480°F.
4.6.4.2 Antiagglomerant Inhibitors
Antiagglomerants were developed out of the necessity to extend the range of subcooling for LDHIs beyond that of KHIs, and AAs can achieve subcooling of greater than 40°F. Unlike KHIs, which delay the formation of hydrates, AAs allow their formation at normal rates, but as small nonagglomerating hydrate crystals that are dispersed into an oil or condensate preventing the formation and accumulation of large hydrate crystals. Thus, AAs are suitable only in the presence of liquid hydrocarbon. The mechanism of dispersion is emulsification with the AAs acting as emulsification agents.
AAs Compared to Methanol or Glycols
The comparisons of AAs are similar for KHIs except AAs achieve greater subcooling.
AA Screening Considerations
Although AAs are applicable under most producing conditions, certain conditions must be considered when evaluating a potential application.
These conditions include water salinity, emulsification and de-mulsification (i.e., separation), pipeline hydraulics, water cuts, material compatibility, water treating, and downstream impacts.
• Some AAs have a maximum salinity criterion that is normally not exceeded with produced water.
• Since AAs are based on dispersing (i.d., emulsifying) polar hydrate crystals in a nonpolar oil or condensate phase (i.e., continuous phase), they may sometimes require a de-emulsifier for oil and water separation. Further, the addition of a heater upstream or heat coil inside a separator may be required to melt the hydrate crystals.
• Since AAs form crystals that are then dispersed in the liquid hydrocarbon phase, careful consideration of the potential impact on viscosity should be considered including steady state flow, shut-in flow and restart conditions.
• An additional consideration for AAs is that the water cuts (i.e., percent water in the liquids) should be less than 50%. Higher water cuts can invert the emulsion (i.e., change the continuous liquid phase from liquid hydrocarbon to water) and make the AA ineffective.
• AAs can impact the performance of some metallurgy and elastomers, so impacts on existing hardware should be reviewed.
• AAs typically partition (i.e., disperse) to the liquid hydrocarbon phase, but low residuals can remain in the produced water, which can impact toxicity test results.
Residual AA concentration in the hydrocarbon liquid phase could possibly impact downstream processes and should be considered in the context of overall contribution to a total feed-stream.
Glycol Injection Considerations
Glycol has a relatively low vapor pressure and thus does not evaporate into the vapor phase as readily as methanol. The solubility of glycol in liquid hydrocarbons is relatively low. For the above reasons, glycol can be more economically recovered, thus reducing the operating expenses below those of methanol systems. The injection part of the system (items 1 to 5 in Figure 4.32) is similar to the methanol injection system. Taking into consideration that, the viscosities of ethylene glycol and its aqueous solutions increase significantly as temperature decreases. Additional equipment in the glycol system is for recovering and reclaiming the glycol. A three-phase separator (6) separates the water and glycol from the hydrocarbon phases. The water–glycol solution in the separator is sent to the reboiler (7) while gas is delivered to the sales line, and the hydrocarbon condensate is dumped to the condensate storage tanks.
In the reboiler, excess water is boiled away from the glycol. The glycol reconcentrated in the reboiler is then available again for injection into the gas stream. Separation of the glycol water phase from the hydrocarbon-liquid field requires a temperature above 700F and a residence time of 10 to 15 minutes.
Nozzle Design (Figure 4-33)
Due to the vapor pressure of glycol, a fine, well-distributed mist is required to obtain adequate mixing with the gas to ensure optimum results, thus spray nozzles are normally used. Nozzle selection is a major consideration in the design of cold separation facilities or plants using glycol injection. Glycol injection normally takes place just upstream of a heat exchanger or chiller where gas is being chilled. Proper nozzle selection will ensure that the glycol spray covers the tube sheet. 100 to 150 psi differential pressure at the nozzle is sufficient to atomize the glycol.
Process stream velocities should be at least 12 ft/s.
Fig. 4-32. Glycol injection system at wellhead.
Fig. 4-33. Schematic of a spray nozzle used in glycol injection.
Glycol Selection General Guideline
The three glycols normally used to prevent the formation of hydrates are:
• Ethylene glycol (EG)
• Diethylene glycol (DEG)
• Triethylene glycol (TEG)
Selection of a glycol depends on the composition of the hydrocarbon flow stream as follows:
1- If glycol is to be injected into a natural gas transmission line where glycol recovery is of less importance than hydrate protection, ethylene glycol is the best choice because it produces the greatest hydrate depression and has the highest vapor pressure of any of the glycols.
2- If glycol is to be injected into a unit where it will contact hydrocarbon liquids, ethylene glycol is preferred because it has the lowest solubility in high molecular weight hydrocarbons.
3- If vaporization losses are severe, either diethylene or triethylene glycol are the best choice because both have a lower vapor pressure. Sometimes diethylene glycol is used if there is a combined loss of both gas vaporization and liquid solubility.
4- The freezing point of the glycol solution must be lower than the lowest temperature expected in the system. In inhibitor service, glycol concentrations are usually maintained at 70 to 75 weight percent because freezing of the glycol is not a problem at this concentration.
5- Reboiler temperature is dependent on the type of glycol and its concentration.
6- Temperature should be maintained at a level equal to the boiling point of the desired solution.
7- Boiling points for the three glycol types are plotted in Figures 4-34, and 4-35.
For example, from Figure 4-34 the reboiler temperature should be set at 240 0F in order to produce about 70 weight percent ethylene glycol solution at atmospheric pressure (N.B. 760 mm, Hg= 1 atm. Pressure.). Thermal degradation can occur if the boiling point of the pure glycol is exceeded; it should therefore be avoided. Glycol losses for the two-phase gas condensate systems are normally estimated at 1 to 2 gallons per 100 barrels of hydrocarbon liquid produced. Vaporization into the gas stream and solution into the hydrocarbon liquid usually cause only a small portion of the total loss.
The most significant causes of glycol losses are leakage and carryover with the hydrocarbon liquid. Loses also occur from vaporization and carry over in the reboiler.
Fig. 4-34. Boiling and condensation points of Ethylene and Di-ethylene glycol.
Fig. 4-35. Boiling and condensation points of Tri-ethylene glycol.
Chemical Injection System
The three parts of an injection system are:
Pump, meter, and control system
Single-Point Chemical Injection
A single pump, meter, and control system service one injection point.
Disadvantages are:
1- Limited turn-down capability and increased life-cycle cost
2- Weight and space increase as injection points increase
Multi-Point Chemical Injection
A shared pump and multiple meter and control devices servicing multiple injection points.
Advantages:
1- Increased turn-down capacity.
2- Per well capital investment decreases as the number of wells increases
3- Injection points are easily added
4- Lower weight and space requirements for higher quantity well applications
Disadvantages
1- Instrumentation intensive, and multiple control loops required
2- Requires variable speed for fixed crank pumps
3- Experiences high-pressure drops from header to recycle line
Metering Pump Types:
Diaphragm Pumps
Advantages
Hermetically sealed, no contamination to atmosphere
Long-life diaphragms typically greater than 2 years continuous duty (20,000 hours)
Long-life of hydraulic plunger seals, typically greater than 2 years continuous duty.
Internal hydraulic relief
Maximum safeguard to environment and personnel safety.
Disadvantages
Higher purchase price
More complex maintenance required
Plunger Pumps
Advantages
Lower purchase price
Less complicated maintenance (easier to understand)
Disadvantages
Plunger packing service life typically less than 2000 hours
Friction between plunger and packing
Comparison of Hydrate Prevention Methods
Overview
The four methods (indirect heaters, methanol injection, glycol injection, and downhole regulators) discussed above are proven safe and reliable.
Evaluation should consider:
1- Development of CAPEX and OPEX (including chemicals and fuel)
2- Space needs (especially in offshore operations) and operating hazards.
Heaters
1- Capital costs and the fuel expense of heaters are relatively large, and it is difficult to maintain a clean, reliable fuel supply to remote heater locations.
2- Indirect heater requires a large amount of space.
3- Fire boxes with proper flame arrestors have minimized the hazards from fired equipment, but they should be bought with strict attention paid to detailed design.
Chemical Injection
Advantages and disadvantages of methanol injection and glycol injection are listed in Table 4-5.
The use of methanol requires only a free-water separator and a suitable means for injection and atomizer, whereas the use of glycol requires a free-water separator plus a gas–liquid separator and a glycol reconcentration unit at the point of recovery downstream.
Downhole Regulators
No routine service is required on downhole regulators, but a wireline service company must be used each time the pressure drop has to be changed and when the regulator is removed. A well with a downhole regulator may require injection of methanol or glycol when it is brought back online after a shut-in until the flow and temperature stabilize. After a well declines to less than allowable production, the downhole regulator will have to be removed, and another form of hydrate prevention may prove necessary.
Downhole regulators do not present special safety hazards, but because work with regulators involves working in the well, losing the well is always a danger.
Inhibitor Advantages Disadvantages
Methanol Relatively low initial cost Minimal equipment
Simple system
High operating cost
Hauling to site necessary
Disposal procedures and precautions for water-methanol mixture
Glycol Usually lower operating cost than methanol when both systems recover chemical
Simple system High initial cost
Hauling to site necessary
Large loss if line breaks
Table 4-5 Methanol and Glycol Injection Comparisons
Summary of Hydrate Prevention Methods:
The methanol injection system is often used for temporary hydrate prevention service in small installations.
Larger installations are favored for indirect heaters or glycol injection systems.
Downhole regulators are most useful in large high-pressure reservoirs in which excess pressure is available and the reservoir pressure is not expected to decline rapidly.
Table 4-6 contains a summary comparison of the above methods.
4.6.3 Hydrate Inhibition with Methanol and Glycols
The amount of chemical inhibitor required to treat the water in order to lower the hydrate formation temperature may be calculated from the Hammerschmidt equation:
ΔT = KWR / M (100-WR) Eq. 4- 14
Where
ΔT is the depression in hydrate formation temperature (0F),
WR is weight percent of inhibitor for water treatment,
K is a constant that depends on the type of inhibitor (table 4-7), and
M is the molecular weight of the inhibitor (table 4-7)
Table. 4-7. M and K values for hydrate inhibitors.
Once the required inhibitor concentration has been calculated, the mass of inhibitor required in the water phase may be calculated from the following equation:
Amount of inhibitor (mI) for water phase for each lb/MMscf
mI = WR X lbH2O/MMscf /(1- WR) Eq. 4-15
Or in general equation
mI = WR X mH2O / (WL − WR) Eq 4-16
mI = Amount of Inhibitor (lb)
mH2O = Amount of water (lb)
WR = weight percent of inhibitor for water treatment,
WL = Lean inhibitor concentration (for example, 80% EG; WL =0.8, 100% methanol; WL=1, etc.)
Fig. 4-36 . Ratio of methanol vapor composition to methanol liquid composition
Determination of Total Inhibitor Required
Total inhibitor required =
Inhibitor required for free water + Inhibitor lost to vapor phase + Inhibitor soluble in condensate
Where inhibitor lost to vapor phase is determined from Figure 4-36. (Methanol lost to the vapor phase), while glycol vaporization losses are generally very small and are typically ignored in calculations.
Inhibitor soluble in the condensate is determined from Figure 4-37 and a value of approximately 0.5% can be used. Solubility of EG in the liquid hydrocarbon phase is extremely small. A solubility of 0.3 lb per 1000 gal. (U.S.) of NGL is often used for design purposes.
Engineering Data Book “GPSA 2004”, recommended that Eq 4-15, should not be used beyond 20-25 wt% for methanol and 60-70 wt% for the glycols. For methanol concentrations up to about 50%, the Nielsen-Bucklin equation (4-17) provides better accuracy:
ΔT 0F =−129.6 ln (zH2O) Eq 4-17
where
z = mole fraction in the liquid phase
zH2O is mole fraction of water
Mole fraction of inhibitor, zI = 1 - zH2O.
WR = (MI X zI ) / (MI X zI + 18 X zH2O) Eq.4-18
MI = Molecular weight of inhibitor
For methanol, WR = (32 X zI )/ (32 X zI + 18 X zH2O). (Examples 4-10, and 4-11 will explain).
Fig. 4-37. Solubility of Methanol in Paraffinic Hydrocarbons vs. Temperature at Various Concentrations
Example 4-10: Determining the Amount of Methanol Required in a Wet Gas Stream
Given:
Flow rate
Gas= 20 MMscfd (Sp.Gr = 0.600)
Condensate = 800 bpd (60 0API/ SP.Gr.=0.739) = 40 bbl/MMscf
Produced Water = 60 Bwpd (Sp.Gr.= 1.03) = 3 bbl/MMscf
Upstream choke
FWHP = 3000 psia (P1)
FWHT = 100 0F (T1)
Down stream choke (cold line)
FWHP = 2000 psia (P2)
FWHT = 60 0F (T2)
Determine: Calculate the total methanol required to prevent hydrates from forming.
Fig. 4-38. Example 4-9.
Solution:
1. The amount of water that will be condensed is determined from McKetta-Wehe (Figure 4-8), assuming the gas is saturated at wellhead conditions.
Water content upstream choke @ 3,000 psia & 100 0F = 32.0 Ib/MMscf
Water content downstream choke @ 2,000 psia & 60 0F = 11.5 lb/MMscf
Water Condensed =(32.0 – 11.5) = 20.5 lb/MMscf
Produced Water = (1.03 X 2.205 X 159) = (361 lb/bbl) X 3 = 1083 lb/MMscf
Total water = (20.5 + 1083) =1103.6 Ib/MMscf
2. From pressure-temperature curve, the hydrate formation temperature is 68 0F (refer to Figure 4-15), or according to equation 4-12, or Equation 4-13 Hydrate formation temperature is 66 0F.
(In this calculation, we will consider the value 68 0F)
The required dewpoint depression then is 68 0F– 60 0F = 8 0F
3. The concentration of methanol required in the liquid water phase from Equation 4-14 is:
8 0F = 2335 WR / 32 (100 – WR)
Rearranging and solving for WR = 9.892% = 0.09892
4. Therefore, from equation 4-15, the estimated methanol required in the liquid water phase is:
mI = 0.09892 x 1103.65 / (1 – 0.09892 ) = 121 lb/MMscf
or we can use equation mI = WR X mH2O / (WL − WR) Eq 4-16
5. From Figure 4-36, the methanol that will flash into the vapor phase at 2000 psia and 60 0F is
= n lbs.methanol/MMscf / WT% methanol in water phase = 1.52
6. Therefore, the methanol in the vapor phase (n) is:
n= 1.52 X 9.892% =15 lb/MMscf
7. A barrel of our condensate weighs:
= [0.739 x 2.205 (lb per liter water) x 159 (liter per barrel)] = 259 lb/bbl
8. Therefore, the approximate amount of methanol soluble in the condensate or liquid hydrocarbon phase (assuming a 0.5% solubility by weight) = 0.005 x 259 x 40 = 52 lb/MMscf
9. Thus, the total amount of methanol required is:
Total = liquid water phase (121) + Vapor phase (15) + Soluble in condensate (52) =188 lb/MMscf
Total = (188 lb/MMscf) x (20 MMscfd) = 3760 lb/day
Note that for gas-condensate wells producing a reasonable or high amount of condensate, the amount of methanol soluble in the condensate is crucial to determining the amount needed.
Approximately 188 lb of methanol must be added so that approximately 121 lb will be dissolved into the water phase.
Since the specific gravity of methanol is 0.791 (at 68 0F), this is equivalent to:
= 188/ [0.791 x 2.205 (lb per liter water) ] = 108 liter/MMscf
= 28.5 gal/ MMscf
For 20 MMscfd, methanol injection will be 570 GPD (13.6 bbl/d).
Example 4-11—100 MMscf/d of natural gas leaves an offshore platform at 100°F and 1200 psia. The gas comes onshore at 40°F and 900 psia. The hydrate temperature of the gas is 65°F. Associated condensate production is 10 bbl/MMscf. The condensate has an API gravity of 50 and a MW of 140. Calculate the amount of methanol and 80 wt% EG inhibitor required to prevent hydrate formation in the pipeline.
Solution Steps:
Methanol
1. Calculate the amount of water condensed per day from McKetta-Wehe (Figure 4-8),
Win =53.0 lb /MMscf
Wout = 9.5 lb/MMscf
ΔW = 43.5 lb/MMscf
Water condensed = (100)( 43.5) = 4350 lb/day
The required dewpoint depression is 65 0F– 40 0F = 25 0F
2. Calculate required methanol inhibitor concentration from Eq 4-14 and 4-17.
ΔT = KWR / M (100-WR) Eq. 4- 14
25 0F = 2335 WR / 32 (100 – WR)
WR = 25% = 0.25
Using equation 4-17
ΔT 0F =−129.6 ln (zH2O) Eq 4-17
zH2O = 0.825 , i.e. zI = (1- 0.725 )= 0.175
From WR = (MI X zI ) / (MI X zI + 18 X zH2O) Eq.4-18
WR = (0.175 X 32) / (0.175 X 32 + 0.825 X 18) = 0.275
WR = 0.275, (use this value in subsequent calculations)
3. Calculate mass rate of inhibitor in water phase
Amount of inhibitor for water phase for each lb/MMscf
mI (lb)= WR X lbH2O/MMscf /(1- WR) Eq. 4-15
= 0.275 X 43.5 /(1-0.275) = 16.5 lb/MMscf = 1650 lb/day
4. Estimate vaporization losses from Fig. 4-36. @ 40°F and 900 psia, losses = 1.05 lb/MMscf
wt% MeOH
daily losses = (1.05)(100)(27.5) = 2890 lb/day
5. Estimate losses to hydrocarbon liquid phase from Fig.4-37. @ 40°F and 27.5 wt% MeOH, approximately = 0.2 mol% lb • mols of condensate per day– “note mole% not weight %”
6. A barrel of our condensate weighs: (API =50, i.e., Sp.Gr = 0.7796), molecular weight =140.
lb.Mole per day = 100 MMscfd X 10 bbl/MMscf X [0.7796 x 2.205 (lb per liter water) x 159 (liter per barrel)] /140 = = 1950 lb−mols/day.
7. lb • mols methanol = (1950)(0.002) = 3.9 lb • mols/day
lb methanol = (3.9)(32) = 125 lb/day
Total methanol injection rate = 1650 + 2890 + 125 = 4665 lb/day
Since the specific gravity of methanol is 0.791 (at 68 0F), this is equivalent to:
4665 / [0.791 x 2.205 (lb per liter water) ] = 2674 liter/day
= 26.74 Liter/ MMscf
Methanol left in the gas phase can be recovered by condensation with the remaining water in downstream chilling processes. Likewise, the methanol in the condensate phase can be recovered by water by downstream water washing.
80 wt% EG
1. Calculate required inhibitor concentration from Eq 4-14.
ΔT = KWR / M (100-WR) Eq. 4- 14
25 0F = 2335 WR / 62 (100 – WR)
W = 40% = 0.4
2. Calculate mass rate of inhibitor in water phase from Eq.4-16.
mI = WR X mH2O / (WL − WR) Eq 4-16
mI =0.4 X 4350 / (0.8 -0.4) = 4350 lb/day
Vaporization and liquid hydrocarbon losses are negligible.
Inhibitor losses represent a significant operating cost and can cause problems in downstream process units. Efficient inhibitor separation should be provided.
4.6.4 Low Dosage Hydrate Inhibitors (LDHIs)
LDHIs can provide significant benefits compared to thermodynamic inhibitors including:
1- Significantly lower inhibitor concentrations and therefore dosage rates. Concentrations range from 0.1 to 1.0 weight percent polymer in the free water phase, whereas alcohols can be as high as 50%
2- Lower inhibitor loss caused by evaporation, particularly compared to methanol
3- Reduced capital expenses through decreased chemical storage and injection rate requirements; and no need for regeneration because the chemicals are not currently recovered.
4- These are especially appropriate for offshore where weight and space are critical to costs
5- Reduced operating expenses in many cases through decreased chemical consumption and delivery frequency
6- Increased production rates, where inhibitor injection capacity or flowline capacity is limited
7- Lower toxicity
4.6.4.1 Kinetic Hydrate Inhibitors
KHIs were designed to inhibit hydrate formation in flowlines, pipelines, and downhole equipment operating within hydrate-forming conditions such as subsea and cold-weather environments. Their unique chemical structure significantly reduces the rate of nucleation and hydrate growth during conditions thermodynamically favorable for hydrate formation, without altering the thermodynamic hydrate formation conditions (i.e., temperature and pressure). This mechanism differs from methanol or glycol, which depress the thermodynamic hydrate formation temperature so that a flowline operates outside hydrate-forming conditions.
KHIs Compared to Methanol or Glycols
KHIs inhibit hydrate formation at a concentration range of 0.1–1.0 weight percent polymer in the free water phase. At the maximum recommended dosage, the current inhibition capabilities are 28°F of subcooling in a gas system and 20°F in an oil system with efforts continuing to expand the region of effectiveness.
For relative comparison, methanol or glycol typically may be required at concentrations ranging 20 to 50 weight percent respectively in the water phase.
KHI Screening Considerations
Although KHIs are applicable under most producing conditions, certain conditions must be considered when evaluating a potential application, which include water salinity, freezing conditions, hold time (i.e., period of effectiveness), water saturation, and high temperature processes.
• At water salinity levels greater than approximately 17% NaCL, the polymer may come out of solution, thereby reducing KHI effectiveness.
• A solution of KHI in water does not provide protection from freezing or icing conditions, neither in the line being treated nor in the KHI storage tank. If ambient temperatures are expected to fall below freezing, the KHI storage volume must be freeze-protected through the use of insulation on the container and piping or addition of antifreeze (typically ethylene glycol) to the KHI solution.
• A solution of KHI cannot be used for melting ice or hydrate plugs. It is recommended to have other strategies, such as a small quantity of ethylene glycol or methanol for remediation purposes in the event of a blockage.
• The KHI delivery system must be capable of providing sufficient dosage to achieve a hold time greater than the water residence time in the piping. Factors to consider include:
— The design basis duration of hydrate forming conditions for an unplanned shut-in.
— The potential for water to pool in low sections of piping (e.g., turn-down hydraulics, flowline profile, pigging frequency, flowline interconnects that are not used continuously) and dead legs.
— The seasonal duration of the cold point temperature below hydrate temperature, if applicable.
• If the gas is undersaturated with respect to water, the water in the KHI solution will evaporate and leave a high viscosity fluid. This can be addressed by using a more dilute KHI solution, or by changing the KHI carrier fluid to ethylene glycol.
• The KHI and water from the KHI solution will form separate phases if the inhibited fluid is above the lower critical solution temperature (LCST) of the KHI solution.
• The KHI polymer suffers degradation effects at temperatures above 480°F.
4.6.4.2 Antiagglomerant Inhibitors
Antiagglomerants were developed out of the necessity to extend the range of subcooling for LDHIs beyond that of KHIs, and AAs can achieve subcooling of greater than 40°F. Unlike KHIs, which delay the formation of hydrates, AAs allow their formation at normal rates, but as small nonagglomerating hydrate crystals that are dispersed into an oil or condensate preventing the formation and accumulation of large hydrate crystals. Thus, AAs are suitable only in the presence of liquid hydrocarbon. The mechanism of dispersion is emulsification with the AAs acting as emulsification agents.
AAs Compared to Methanol or Glycols
The comparisons of AAs are similar for KHIs except AAs achieve greater subcooling.
AA Screening Considerations
Although AAs are applicable under most producing conditions, certain conditions must be considered when evaluating a potential application.
These conditions include water salinity, emulsification and de-mulsification (i.e., separation), pipeline hydraulics, water cuts, material compatibility, water treating, and downstream impacts.
• Some AAs have a maximum salinity criterion that is normally not exceeded with produced water.
• Since AAs are based on dispersing (i.d., emulsifying) polar hydrate crystals in a nonpolar oil or condensate phase (i.e., continuous phase), they may sometimes require a de-emulsifier for oil and water separation. Further, the addition of a heater upstream or heat coil inside a separator may be required to melt the hydrate crystals.
• Since AAs form crystals that are then dispersed in the liquid hydrocarbon phase, careful consideration of the potential impact on viscosity should be considered including steady state flow, shut-in flow and restart conditions.
• An additional consideration for AAs is that the water cuts (i.e., percent water in the liquids) should be less than 50%. Higher water cuts can invert the emulsion (i.e., change the continuous liquid phase from liquid hydrocarbon to water) and make the AA ineffective.
• AAs can impact the performance of some metallurgy and elastomers, so impacts on existing hardware should be reviewed.
• AAs typically partition (i.e., disperse) to the liquid hydrocarbon phase, but low residuals can remain in the produced water, which can impact toxicity test results.
Residual AA concentration in the hydrocarbon liquid phase could possibly impact downstream processes and should be considered in the context of overall contribution to a total feed-stream.
Fundamentals of Oil and Gas Processing
Basics of Gas Field Processing
Basics of Corrosion in Oil and Gas Industry
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Basics of Gas Field Processing
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Re: Basics of Gas Field Processing Book "Full text"
Gas Dehydration - Chapter 5 - Part 1
Chapter 5 113
Gas Dehydration 113
Introduction 113
Absorption 114
5.1. Gas Dehydration by Absorption 114
5.1.1 Absorption and Stripping 114
5.1.2 Raoult and Dalton’s Laws 114
5.1.3 Glycol-Water Equilibrium 115
5.2 Glycol Dehydration 115
5.2.1 Principles of Operation 115
5.3 Effect of Operating Variables 123
5.3.1 Glycol Selection 123
5.3.2 Inlet Gas Temperature 124
5.3.3 Lean Glycol Temperature 124
5.3.4 Glycol Reconcentrator Temperature 124
5.3.5 Temperature at Top of Still Column 125
5.3.6 Contactor Pressure 125
5.3.7 Reconcentrator Pressure 125
5.3.8 Glycol Concentration 126
5.3.9 Glycol Circulation Rate 128
5.3.10 Number of Absorber Trays 129
5.4 Enhanced Glycol Concentration Processes 132
5.4.2 DRIZO® (wt.-2) Process 133
5.4.3 Coldfinger® Process 135
5.5 Other Considerations of Glycol Dehydration 136
5.5.1 Systems Utilizing Glycol-Gas Powered Pumps 138
5.5.2 Systems Utilizing Electric Driven Pumps 139
5.6 Glycol Gas Contactors 140
5.6.1 Trays and Packing 140
5.7 Glycol Dehydration System Sizing 146
5.7.1 Contactor Diameter 146
5.7.2 Number of Trays and Tray Spacing 148
5.7.3 Downcomers 148
5.7.4 Glycol Circulation Rate 149
5.7.5 Lean Glycol Concentration 149
5.7.6 Glycol-Glycol Preheater 149
5.7.7 Glycol-Gas Cooler 149
5.7.8 Glycol-Glycol Heat Exchanger 149
5.7.9 Gas-Glycol-Condensate Separator 150
5.7.10 Reconcentrator 150
5.7.11 Heat Duty 150
5.7.12 Fire Tube Sizing 151
5.7.13 Reflux Condenser 154
5.7.14 Stripping Still Column 154
5.7.15 Filters 155
5.8 Calculation Examples for Glycol Dehydration 158
Example 5-1 158
Example 5-2 159
Example 5-3 159
5.9 Glycol Unit Operation 161
5.9.1 Start up 161
5.9.2 Routing Operation 161
5.9.3 Shut Down 161
5.10 Glycol Maintenance and Care 162
5.10.1 Preventive Maintenance 162
5.11 Glycol Operation Considerations 165
5.11.1 Oxidation 165
5.11.2 Thermal Decomposition 165
5.11.3 pH Control 166
5.11.4 Salt Contamination and Deposits 166
5.11.5 Hydrocarbons 166
5.11.6 Sludge 167
5.11.7 Foaming and defoamers 167
5.12 Analysis and Control of Glycol 167
5.12.1 Visual Inspection 167
5.12.2 Chemical Analysis 167
5.12.3 Chemical Analysis Interpretation 168
5.13 Troubleshooting 5.13.1 General Considerations 172
5.13.2 Main approach to troubleshooting: 172
5.13.3 High Dew Points 172
5.13.3 Glycol Loss from the Contactor 173
5.13.5 Glycol Loss from the Reconcentrator 174
5.3.16 Glycol Loss From Glycol Hydrocarbon Separator 175
5.13.7 Glycol Loss—Miscellaneous 175
5.14 Glycol System Cleaning 175
5.15 Eliminating Operating Problems 176
5.15.1 Inlet Scrubber/Microfiber Filter Separator 176
5.15.2 Absorber 177
5.15.3 Glycol-Gas Heat Exchanger 178
5.15.4 Lean Glycol Storage Tank or Accumulator 178
5.15.5 Stripper or Still Column 179
5.15.6 Reboiler 180
5.15.7 Stripping Gas 181
5.15.8 Glycol Circulating Pump 182
5.15.9 Flash Tank or Glycol-Gas Separator 183
5.15.10 Gas Blanket 183
5.15.11 Reclaimer 184
5.16 Improving Glycol Filtration 184
Adsorption 185
5.17 Overview of Adsorption Processes 185
5.18 Properties of Industrial Adsorbents for Dehydration 187
5.19 Solid Bed Adsorption Process 188
5.20 Principles of Operation 188
5.20.1 Introduction 188
5.20.2 Drying Cycle 189
5.20.4 Regeneration Cycle 189
5.21 Adsorption System Performance 190
5.22 Effect of Process Variables 191
5.22.1 Quality of Inlet Gas 191
5.22.2 Temperature 191
5.22.3 Pressure 192
5.22.4 Cycle Time 192
5.22.5 Gas Velocities 192
5.22.6 Source of Regeneration Gas 192
5.22.7 Direction of Gas Flow 193
5.22.8 Desiccant Selection 194
5.22.9 Effect of Regeneration Gas on Outlet Gas Quality 195
5.22.10 Pressure Drop Considerations 196
5.22.11 Equipment 197
5.23 Desiccant Performance 200
5.24 Design 202
5.24.1 Pressure Drop & Minimum Diameter 202
5.24.2 Mass desiccant Required & Bed Length 203
5.24.3 Regeneration Calculations 205
5.24.4 Design Example 208
5.25 Nonregenerable Dehydrator 211
5.25.1 Calcium Chloride Dehydrator Unit 211
5.25.2 Principles of Operation 211
5.26 Dehydration by Refrigeration 213
5.27 Dehydration by Membrane Permeation 214
5.28 Other Processes 214
5.29 Comparison of Dehydration Processes 215
-----------------
Chapter 5
Gas Dehydration
Introduction
After the liquid (free) water has been removed from the gas stream by separation, 25 to 120 lbs of water per MMscf of gas will remain, depending on the temperature and pressure of the gas.
The warmer the inlet gas and the lower the pressure, the more water vapor the gas stream will contain.
Dehydration is the process used to remove water from natural gas and natural gas liquids (NGLs), and is required to:
1- prevent formation of hydrates and condensation of free water in processing and transportation facilities,
2- meet a water content specification, and
3- prevent corrosion
Techniques for dehydrating natural gas, associated gas condensate and NGLs include:
1- Absorption using liquid desiccants,
2- Adsorption using solid desiccants,
3- Dehydration with CaCl2,
4- Dehydration by refrigeration,
5- Dehydration by membrane permeation,
6- Liquid dehydration by gas stripping, and
7- Liquid dehydration by distillation
Fig.5-1. Dehydration of natural gas.
Absorption
5.1. Gas Dehydration by Absorption
5.1.1 Absorption and Stripping
In the absorption process, a hygroscopic liquid is used to contact wet gas and remove the water vapor. Through absorption, the water in a gas stream is dissolved in a relatively pure liquid solvent stream. Normally, between 20 to 115 lbs of water per MMscf of gas must be removed before the required dew point of the gas is met. The reverse process, in which the water in the solvent is transferred into the gas phase, is known as stripping. The terms regeneration, re-concentration, and reclaiming are also used to describe stripping (or purification) because the solvent is recovered for reuse in the absorption step. Absorption and stripping are frequently used in gas processing, gas sweetening, and glycol dehydration.
5.1.2 Raoult and Dalton’s Laws
Absorption can be qualitatively modeled by using Raoult’s and Dalton’s laws.
For a vapor liquid equilibrium system:
Raoult’s Law state that the partial pressure of a component in a vapor phase that is in equilibrium with a liquid is directly proportional to the mole fraction of the component in the liquid phase. Or, the partial pressure of a component in a vapor phase that is in equilibrium with a liquid equal the vapor pressure of its pure component multiply by mole fraction of the component in liquid phase.
Dalton’s Law states that the partial vapor pressure of a component is equal to the total pressure multiplied by its mole fraction in the gas mixture.
Raoult’s Law expressed in equation form is:
pi = PiXi Eq. 5-1
Dalton’s Law expressed in equation form is:
pi = PYi Eq. 5-2
Where:
pi = Partial vapor pressure of component i
Pi = Vapor pressure of pure component i
Xi = Mole fraction of component i in the liquid
P = Total pressure of the gas mixture
Yi = Mole fraction of component i in the vapor
Combining these laws we have:
PYi = piXi
or
pi/P = Yi/Xi Eq. 5-3
Since the pure-component vapor pressure and the total pressure are not affected by composition.
Equation (5-3) is significant. It states that the ratio of the vapor mole fraction to the liquid mole fraction for any component is independent of the concentrations of that component and the other components present. The ratio Yi /Xi is commonly known as the K-value.
Since the pure component vapor pressure increases with temperature, the K-value increases with increasing temperature and decreases with increasing pressure.
In physical terms this means:
• Transfer from the gas phase to the liquid phase (absorption) is more favorable at lower temperature and high pressures.
• Transfer to the gas phase (stripping) is more favorable at higher temperatures and lower pressures.
5.1.3 Glycol-Water Equilibrium
Absorption processes are dynamic and continuous.
Gas flow cannot be stopped to let the vapor and liquid reach equilibrium. Thus, the system must be designed to approach equilibrium as closely as possible while flow continues. This is accomplished by using a trayed or packed contactor in which the gas and liquid are in counter current flow. The closer to 100% equilibrium that a tray or packed section approaches, the higher the tray or packing efficiency.
For example, A common tray efficiency is 25%, meaning that 25% of the water molecules that would have been transferred under equilibrium conditions were actually transferred.
Wet gas enters the bottom of the column and contacts the rich glycol (high water content) just before the glycol leaves the column. The gas encounters leaner glycol as it works its way up the column, contacting the leanest glycol (lowest water content) just before it leaves the column.
The equilibrium based on Dalton’s and Raoult’s Laws can be rearranged as follows:
Yi = Xi (Pi / P) Eq. 5-4
Since Pi/P is constant for constant temperature, the concentration of the water in the gas must be directly proportional to the concentration in the liquid. However, the liquid concentration is constantly changing as water is absorbed. The counter current flow in the contactor makes it possible for the gas to transfer a significant amount of water to the glycol and still approach equilibrium with the leanest glycol concentration.
5.2 Glycol Dehydration
5.2.1 Principles of Operation
The most common liquid used in absorption type dehydration units is triethylene glycol (TEG)
Liquid desiccant dehydration equipment is simple to operate and maintain. It can easily be automated for unattended operation; for example, glycol dehydration at a remote production well. Liquid desiccants can be used for sour gases, but additional precautions in the design are needed due to the solubility of the acid gases in the desiccant solution. At very high acid gas content and relatively higher pressures the glycols can also be “soluble” in the gas.
Glycols are typically used for applications where dew point depressions of the order of 60° to 120°F are required.
Diethylene glycol (DEG), triethylene glycol (TEG), and tetraethylene glycol (TTEG) are used as liquid desiccants, but TEG is the most common for natural gas dehydration. The schematics in Figures 5-2 and 5-3 show the flow through a typical glycol dehydration system. The glycol dehydration process can be discussed in two parts gas system (Figure 5-2) & glycol system (Figure 5-3)
1. Gas System (figure 5-2)
Wet gas enters the unit through the inlet gas scrubber/microfiber filter separator, usually vertical, to remove liquid and solid impurities. After passing through the microfiber filter separator, the gas enters the glycol gas contactor near the bottom of the vessel.
Even if the dehydrator is near a production separator. The inlet gas scrubber upstream the contactor (or filter) will prevent accidental dumping of large quantities of water (fresh or salty), hydrocarbons, treating chemicals or corrosion inhibitors into the glycol contactor. Even small quantities of these materials can result in excessive glycol losses due to foaming, reduced efficiency, and increased maintenance. Integral separators at the bottom of the contactor are common. Figure 5-4.
Fig. 5-2. Gas system
The inside of the contactor contains either packing or several trays with weirs that maintain a specific level of glycol so that the gas must bubble through the glycol as the gas flows up.Different types of trays will be presented.
As the wet gas passes upward through each succeeding tray, it gives up the water vapor to the glycol and becomes progressively drier. Before leaving the contactor the gas passes through a mist extractor to remove glycol that may be trying to leave the gas.
Dry gas exits the contactor at the top and passes through an external glycol gas heat exchanger where it cools the incoming dry glycol to increase its absorption capacity (Figure 5-2).
Some installations incorporate a glycol knockout drum (centrifugal separator) which recovers any glycol that has escaped with the gas through the mist extractor.
The dry gas then leaves the dehydrator unit.
Fig. 5-4. Integral separators at the bottom of the contactor
2- Glycol System (figure 5-3)
The glycol pump pumps up dry concentrated glycol to contactor, and then passes through the glycol gas heat exchanger before entering the contractor tower.
The glycol gas heat exchanger cools the glycol to near the temperature of the gas before the glycol enters the contactor.
It is important that the glycol be near the gas temperature to prevent gas from exceeding equilibrium temperature, and to prevent foaming. Dry glycol from the glycol gas heat exchanger enters the contactor tower and flows across the top tray.
This is the first contact between the glycol and gas. Glycol flows downward through downcomers in the tower, absorbing more water as it passes across each tray.
The downcomer seals the glycol passage into the tray below, thus preventing gas from short-circuiting past the bubble caps.
As the glycol flows downward through each succeeding tray, it becomes wetter with the water it has absorbed from the gas and collects in the bottom of the contactor saturated with water.
As the gas moves upward through each succeeding tray, it becomes drier.
The wet glycol that has accumulated in the bottom of the contactor passes through a strainer (filter), which removes abrasive particles, before flowing through the power side of the glycol pump (energy exchange pumps), where it furnishes the power to pump the dry glycol into the contactor. Power comes from the increased head caused by the absorbed water contained in the rich glycol.
From the glycol gas contactor the cool wet glycol passes through a coil (reflux condenser) in the top of the reboiler still column. The coil cools the vapors leaving the still column and condenses the glycol vapors to liquid.
The glycol liquid droplets gravitate back down the still column to the re-concentrator. The water remains as a vapor and continues on out the top of the still column. The cooling coil is commonly called the reflux condenser.
The slightly warmed wet glycol leaving the reflux condenser passes through the glycol-glycol preheater. The hot dry glycol from the glycol reconcentrator heats the wet glycol further, and in turn further cools the dry glycol before it goes to the glycol pumps.
After leaving the glycol-glycol preheater, the heated wet glycol is sent to a low-pressure gas-glycol-condensate separator, where most of the entrained gas and liquid hydrocarbons that were picked up by the glycol on its path through the contactor are removed.
The heat provided by the glycol-glycol preheater helps in the separation of hydrocarbons from the wet glycol. The hydrocarbon condensate is separated from the glycol by a three-phase gas-glycol-condensate separator (Fig. 5-5), or a vertical three-phase separator as in fig.5-3.
Fig. 5-5. Gas-glycol-condensate separator
After the gas and condensate has been separated in the gas-glycol-condensate separator, the wet glycol passes through a microfiber filter (Fig. 5-6). These filters are used to remove solids, tarry hydrocarbons, or other impurities.
Fig. 5-6. Microfiber filter.
From the microfiber filter the wet glycol enters a charcoal or carbon filter. Activated carbon granules in this filter absorb liquid-entrained hydrocarbons, well-treating chemicals, compressor oils, and other impurities that may cause foaming.
From the charcoal filter, the wet glycol flows through the dry glycol to the wet glycol heat exchanger. This heat exchanger preheats the wet glycol as much as possible before entering the glycol reconcentrator, thus reducing the heat duty of the glycol reconcentrator.
From the glycol/glycol heat exchanger, the wet glycol enters the still column which sits vertically atop the glycol reconcentrator (Fig. 5-7). The inside of the still column is packed with either ceramic saddles or stainless steel pall rings, which are used to add surface area and distribute heat to the incoming glycol.
Fig. 5-7. Still column
The incoming wet glycol spreads out uniformly and drips down through the packed section.
The vapors traveling upward from the glycol reconcentrator heats the packing. As the glycol travels down through the heated packing, water begins to be driven off as steam. Units utilizing efficient heat exchangers may remove as much as 75 to 80% of the water contained in the glycol in the still column before the glycol reaches the reconcentrator.
As water vapor travels up through the still column and exits from the top, it carries with it trapped glycol vapor. To prevent the loss of glycol vapor, the still column utilizes a “reflux condenser” located on the top of the packed still column.
Glycol vapors escaping the still column with the steam are attracted to the film of condensed liquid (primarily water) covering the coil surface area where they too are condensed. The liquid droplets gravitate back down the still column into the reconcentrator for further treating, thus preventing excessive glycol loss due to vaporization.
On some units, the glycol enters the still column below the packed section of the column.
Vaporization takes place in the reconcentrator. The reflux condenser operates the same in both types of still columns. Condensed liquid from the reflux condenser drops back into the packed section providing a liquid film over the upper portion of the packing. Glycol vapors escaping with steam from the reconcentrator must pass through the packed section.
The watery film covering the packing recaptures the glycol vapor, condensing it into droplets, which wash back into the reconcentrator. Thus, more glycol vapor can be recovered in this configuration than in the previously described still column.
Since vaporization occurs primarily in the reconcentrator, the operating temperature is lower in this type of still column. This translates into:
- Greater reflux condensation
- Requires larger heat duty
From the packed still column, the wet glycol drops downward into the reconcentrator. The glycol is heated to a temperature at which most of the remaining water and some of the glycol are vaporized. A heat source heats the glycol to between 350 0Fand 400 0F, where:
- It removes the remaining water.
- It is below the decomposition point of TEG.
The temperature of the glycol in the reconcentrator is critical and must be controlled at this point.
Sources of heat include:
- Direct fired (natural draft/forced draft)
- Waste heat (exhaust gases from compressors or generators)
- Electric heaters
The heated vapor (both glycol and water) rises upward through the still column.
As the mixture passes the cool reflux condenser coils, the glycol vapors condense and drop back down. The water vapor leaves the top of the still column as steam.
Some of the steam will condense, so a downspout is provided to drain the water off.
A weir maintains a level of glycol over the heat source, which:
- Prevents over heating of the tubes
- Prevents premature tube failure
As the glycol is purified, it spills over the weir into a separate compartment. From the reconcentrator, the dry (lean) glycol flows to the accumulator surge tank, where the glycol pump raises it to contactor pressure to start another cycle.
Figure 5.8 shows a typical, simplified flow sheet for a glycol absorption unit. The wet gas passes through an inlet scrubber to remove solids and free liquids, and then enters the bottom of the glycol contactor. Gas flows upward in the contactor, while lean glycol solution (glycol with little or no water) flows down over the trays. Rich glycol absorbs water and leaves at the bottom of the column while dry gas exits at the top. The rich glycol flows through a heat exchanger at the top of the still where it is heated and provides the coolant for the still condenser. Then the warm solution goes to a flash tank, where dissolved gas is removed. The rich glycol from the flash tank is further heated by heat exchange with the still bottoms, and then becomes the feed to the still. The still produces water at the top and a lean glycol at the bottom, which goes to a surge tank before being returned to the contactor.
Fig. 5-8. Schematic of typical glycol dehydrator unit.
Fig. 5-9. Glycol dehydration system
5.3 Effect of Operating Variables
Several operating and design variables have an important effect on the successful operation of a glycol dehydration system.
5.3.1 Glycol Selection
Glycols are the most commonly used liquid desiccants in the absorption process because they are:
1. Highly hygroscopic (readily absorb and retain water)
2. Stable to heat and chemical decomposition at the temperature and pressures necessary in the process Low vapor pressures, which minimize equilibrium loss of the glycol in the residual natural gas stream and in the regeneration systebm
3. Easily regenerated (water removed) for reuse
4. Noncorrosive and nonfoaming at normal conditions; impurities in the gas stream can change this, but even then inhibitors can help to minimize these problems
5. Readily available at moderate cost
Hygroscopicity of glycols is affected by the concentration (glycol-to-water ratio), that is, increasing as the concentration increases.
Dew point depression obtainable in a gas stream increases as the glycol concentration increases.
Ethylene Glycol (EG)
Ethylene Glycol tends to have high vapor losses to gas when used in a contactor. It is used as a hydrate inhibitor where it can be recovered from the gas by separation at temperatures <50 0F.
Diethylene Glycol (DEG)
DEG (Diethylene) is somewhat cheaper to buy and sometimes is used for this reason. But, by the time it is handled and added to the units there is no real saving. Compared to TEG, DEG has a larger carry-over loss, offers less dewpoint depression and regeneration to high concentrations is more difficult. For these reasons, it is difficult to justify a DEG unit, although a few units are built each year. Diethylene Glycol reconcentrates at temperatures between 3150 and 325 0F, which yields purity of 97.0%. It degrades at 3280F. It cannot achieve the concentration required for most applications.
Triethylene Glycol (TEG)
Triethylene Glycol is most commonly used in glycol dehydration. It reconcentrates at temperatures between 3500 and 400 0F, which yields purity of 98.8%.
It degrades at 404 0F. It tends to experience high vapor losses to gas at temperatures in excess of 1200F. With stripping gas, dew point depressions up to 1500F are possible.
TEG (Triethylene) is preferred for use in dehydration units because:
• It is more easily regenerated due to its high boiling point and other physical properties.
• It has lower vaporization losses than other glycols
• It has lower capital and operating costs than other glycol systems
Tetraethylene Glycol (TTEG)
TTEG (Tetraethylen) is more viscous and more expensive than the other processes. It reconcentrates at temperatures between 4000 and 4300F. It experiences lower vapor losses to gas at high gas contactor temperatures. It degrades at 4600F. The Only real advantage is its lower vapor pressure which reduces absorber carry-over loss. It may be used in those relatively rare cases where glycol dehydration will be employed on a gas whose temperature exceeds about 50°C [122°F].
Table 5-1. Physical properties of glycols and methanol.
5.3.2 Inlet Gas Temperature
At constant pressure, the water content of the inlet gas increases as the temperature increases.
For example, at 1000 psia and 800F, gas holds 34 lb of water/MMscf. While, at 1200F, gas holds 104 lb of water/MMscf.
If the gas is saturated at the higher temperature, the glycol will have to remove about three times as much water to meet the specifications.
Temperatures above 1150F result in high glycol losses, thus requires tetraethylene glycol.
Temperature should not fall below the hydrate formation temperature range (650 to 700F) and always-above 500F.
Temperatures below 500F cause problems due to an increase in glycol viscosity.
Temperatures below 600 to 700F can cause a stable emulsion with liquid hydrocarbons in the gas and cause foaming in the contactor.
An increase in gas temperature increases the gas volume, which in turn increases the diameter of the glycol contactor.
5.3.3 Lean Glycol Temperature
Dry glycol temperatures entering the top tray of the contactor (approach temperature) should be held low (100 to 150F) above the inlet gas temperature.
Equilibrium conditions between the glycol and the water vapor in the gas is affected by temperature. Glycol entering the top tray of the contactor may raise the temperature of the gas surrounding it and prevent the gas giving up its remaining water vapor.
Inlet glycol temperatures greater than 150F above the gas temperature results in high glycol losses to the gas. Drastic temperature differential also has a tendency to emulsify the glycol with any contaminants subsequent glycol loss.
5.3.4 Glycol Reconcentrator Temperature
Reconcentrator temperature controls the concentration of the water in the glycol.
With a constant pressure, the glycol concentration increases with higher reconcentrator temperature. Reconcentrator temperature should be limited to between 3500 and 4000F.
Minimizes degradation of TEG which begins to degrade at 4040F. At a 400°F (204°C), the typical maximum regeneration temperature, TEG yields a lean-glycol concentration of 98.6 wt% at sea level. Higher purity requires reduction of the partial pressure of water in the reboiler vapor space. The most common way to achieve this pressure reduction is to use a stripping gas, or vacuum distillation, which yields lean glycol concentrations of 99.95 wt% and 99.98 wt%, respectively (Detailed process description will be followed in “Enhanced Glycol Concentration Processes” section).
5.3.5 Temperature at Top of Still Column
A high temperature in the top of the still column can increase glycol losses due to excessive vaporization. A reboiler temperature in the range of 3500 to 4000F insures adequate heat transfer to the ceramic packing in the still column.
The still column operates best (allows the steam to escape) when the vapor outlet temperature is between 2150 and 2250F.
When the temperature reaches 2500F and above, glycol vaporization losses increase.
Still top temperature can be lowered by increasing the amount of glycol flowing through the reflux condenser coil. If the temperature in the top of the still column drops too low, (below 2200F) too much water can be condensed and washed back into the reconcentrator, which increases the reconcentrator heat duty. Too much cool glycol circulation in the reflux condenser coil can lower the still top temperature below 2200F, which can cause the excess water to condense. Thus, most reflux condenser coils have a bypass valve, which allows manual or automatic control of the stripping still temperature.
5.3.6 Contactor Pressure
At a constant temperature, the water content of the inlet gas decreases with an increase in pressure. The lower the pressure, the larger the contactor diameter required.
Good dehydration can be achieved at any pressure below 3000 psig as long as the pressure is constant. Optimum dehydration pressure is often in the range of 550 to 1200 psig.
Sizing calculations should always be based on minimum expected operating gas pressure.
Rapid pressure changes translate into rapid velocity changes in the contactor which:
- Breaks the liquid seals between the downcomers and the trays
- Allows the gas to channel up through both the downcomer and bubble caps
- Allows the glycol to be swept out with the gas
5.3.7 Reconcentrator Pressure
Reducing the pressure in the reconcentrator at a constant temperature results in higher glycol purity. Most reconcentrators operate between 0.25 to 0.75 psig of pressure.
On standard atmospheric reconcentrators, pressures in excess of 1 psig results in:
- Glycol loss from the still column
- Reduction of lean glycol concentration
- Reduction in dehydration efficiency
Pressures more than 1 psig are usually associated with excess water in the glycol and create a vapor velocity exiting the still great enough to sweep glycol out.
Fouled still column packing often contributes to high reconcentrator pressure.
Still column should be adequately vented and packing replaced periodically to prevent backpressure on the reconcentrator.
Pressures below atmospheric will increase the lean glycol concentration because the boiling temperature of the rich glycol/water mixture decreases.(will be discussed later)
Reconcentrators are rarely operated in a vacuum due to the added complexity and the fact that air leaks will result in glycol degradation.
If lean glycol concentrations in the range of 99.5% are required consider:
Operating the reconcentrator at a pressure 500 mm Hg absolute (10 psia), or Using stripping gas Fig. 5-9 can be used to estimate the effect of operating in a vacuum on lean glycol concentration.
Fig. 5-9. Glycol purity versus reconcentrator temperature at different levels of vacuum.
5.3.8 Glycol Concentration
The water content of the dehydrated gas depends primarily on the lean glycol concentration.
The higher the concentration of lean glycol entering the contactor, the greater the dew point depression for a given circulation rate and number of trays. Increasing the glycol concentration above a 99% purity can lead to dramatic results on the outlet dew point (Fig. 5-10).
For example, with a 100 0F inlet gas temperature (110 0F top tray temperature), an outlet dew point of
10 0F can be obtained with 99.0% TEG,
-30 0F can be obtained with 99.8% TEG
-40 0F can be obtained with 99.9% TEG
Higher concentrations of TEG can be obtained by:
- Increasing the glycol reconcentration temperature
- Injecting stripping gas into the reconcentrator
- Reducing the operating pressure of the reconcentrator
Fig.5-10. Equilibrium water dew points with various concentrations of TEG.
Reconcentration temperatures for TEG normally run between 380 0F and 400 0F, which results in glycol purities of 98% to 99%. Figures 5-11 and 5-12 illustrate the effect of stripping gas.
If gas is injected directly into the reconcentrator (via a sparger tube), the concentration of TEG increases significantly from 99.1% to near 99.6% as the gas rate is increased from 0 to 4 SCF/gal.
When the Stahl method is used (counter current gas stripping after the reconcentrator – will be illustrated here later), concentrations as high as 99.95% TEG can be attained at a 400 0F reconcentrator temperature.
Fig. 5-11. Effect of stripping gas on TEG concentration.
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Fig. 5.12. Effect of stripping gas on the concentration using Stahl column
5.3.9 Glycol Circulation Rate
When the number of absorber trays and lean glycol concentration are held constant, the dew point depression of a saturated gas is increased as the glycol circulation rate is increased.
The more lean glycol that comes into contact with the gas, the more water vapor is absorbed out of the gas. Whereas the glycol concentration mainly affects the dew point of the dry gas, the glycol rate controls the total amount of water that can be removed. The normal operating level in a standard dehydrator is 3 gallons of glycol per pound of water removed (Range 2-7).
Figure 5-13 shows that a greater dew point depression is a factor of both glycol concentration and glycol circulation rate.
Fig. 5-13. Calculated dew point depression versus circulation rate (1 equilibrium tray (4 actual trays)).
Excessive circulation rates:
- Overload the reconcentrator and/or decrease the reconcentrator temperature
- Prevent good glycol regeneration, resulting in decrease lean glycol concentration.
- Prevent an adequate glycol gas contact in the absorber; this may results in a decrease of water removed by glycol from the gas.
- Increase pump maintenance problems
- Increase glycol loses
- Only if the reconcentrator temperature remains constant, an increase in circulation rate will lower the dew point of the gas.
5.3.10 Number of Absorber Trays
When the glycol circulation rate and the lean glycol concentration are held constant, the dew point depression of a saturated gas is increased as the number of trays is increased.
Actual trays do not reach equilibrium, and their approach to it is expressed as a fraction of a theoretical tray. A tray efficiency of 25% is commonly used for design.
Four actual trays with efficiencies of 25% would accomplish the job of one theoretical tray.
The number of actual trays in a design ranges from 4 to 12.
An approximation of the actual number of valve trays per foot of packing can be obtained from
Figure 5-14.
Fig. 5-14 Trays of packing required for glycol dehydrator.
For high performance units, the specification of more than 4 trays in a new design can achieve fuel savings (for the same dew point depression) due to:
- Lower circulate rate
- Lower reconcentration temperature
- Lower stripping gas rate
Figure 5-15 shows that specifying a few additional trays in the contactor is much more effective than increasing the glycol circulation rate. The additional investment for a taller contactor is often justified by fuel savings.
Evaluation of a TEG system involves first establishing the minimum TEG concentration required to meet the outlet gas water dewpoint specification. Fig. 5-10 shows the water dewpoint of a natural gas stream in equilibrium with a TEG solution at various temperatures and TEG concentrations. Fig. 5-10 can be used to estimate the required TEG concentration for a particular application or the theoretical dewpoint depression for a given TEG concentration and contactor temperature. Actual outlet dewpoints depend on the TEG circulation rate and number of equilibrium stages, but typical approaches to equilibrium are 10-20°F. Equilibrium dewpoints are relatively insensitive to pressure and Fig. 5-10 may be used up to 1500 psia with little error.
Fig. 5-15 Effect of number of absorber trays on dew point depression.
When dehydrating to very low dewpoints, such as those required upstream of a refrigeration process, the TEG concentration must be sufficient to dry the gas to the hydrate dewpoint.
Once the lean TEG concentration has been established, the TEG circulation rate and number of trays (height of packing) must be determined. Most economical designs employ circulation rates of about 2-5 gal. TEG/lb H2O absorbed. The relationship between circulation rate and number of equilibrium stages (N) employs the absorption calculation techniques set out in Engineering Data Book “GPSA 2004” Chapter 19. Calculation results for TEG systems are presented in figures 5-36 through 5-41 where they show the fraction of water removed versus TEG rate with respect to different glycol purities. (N=1 theoretical trays, 4 actual trays)
The graphs in these figures apply only if the feed gas is water saturated. They are based on feed-gas and therefore contactor temperatures of 80°F, but are essentially independent of temperature.
Although the K values in the absorption factors (i.e., L/VK) do increase with temperature, the required TEG rates also increase, and this tends to compensate for the increasing K values and keep the absorption factors fairly constant.
Conversion from equilibrium stages to actual trays can be made assuming an overall tray efficiency of 25-30%. For random and structured packing, Height of Packing Equivalent to a Theoretical Plate (HETP) varies with TEG circulation rate, gas rate, gas density, and packing characteristics but a value of about 60 inches is usually adequate for planning purposes.
When the gas density exceeds about 6 lb/ft3 (generally at very high pressures), the above conversions may not provide sufficient packing height and number of trays. Also when the contactor temperature is less than about 60°F, the increased viscosity of the TEG can reduce mass transfer efficiency, and temperatures in this range should be avoided.
Typical tray spacing in TEG contactors is 24 inches. Therefore, the total height of the contactor column will be based on the number of trays or packing required plus an additional 6-10 ft to allow space for vapor disengagement above the top tray, inlet gas distribution below the bottom tray, and rich glycol surge volume at the bottom of the column. Bubble cap trays have historically been used in glycol contactors due to the low liquid rates versus gas flow, but structured packing is widely used. Generally, structured packing allows a significantly smaller contactor diameter and a slightly smaller contactor height. Calculation examples will be followed.
5.4 Enhanced Glycol Concentration Processes
The enhanced glycol concentration processes include stripping gas, vacuum, Drizo, and cold finger.
5.4.1 Stripping Gas and Vacuum
Normal dehydration systems with TEG glycol purity of 98.5% are capable of achieving dew point depressions up to 70 0F.
If very pure glycol (up to 99.9% TEG) is required and cannot be achieved by the standard regeneration system, stripping gas may be used. A small amount of dry natural gas, normally taken from the fuel stream, is injected into the reconcentrator. Since hot gas has an affinity for water, the stripping gas is bubbled through the hot glycol, which strips the remaining water from the glycol.
This gas can be put directly into the reconcentrator or it can be added to the storage tank where it can percolate through the packed column between the two vessels (Stahl column) fig 5-17.
The Stahl column also serves as a weir where the dry glycol spills downward by gravity over packing while the gas goes upward, removing even more water.
This method prevents air from coming into contact with the dry glycol in the storage tank, thus preventing oxidization of the glycol.
Oxygen entry into the glycol system will:
- Decompose the glycol to some extent
- Cause corrosion within the system.
Stripping gas can:
- Reduce the temperature at which the reconcentrator must operate
- Reduce the glycol circulation rate necessary to dehydrate the gas adequately
Stripping gas may be used to obtain higher dew point depressions.
Vacuum-operated glycol units can achieve glycol purities of up to 99.9% but are rarely used because of:
- High operating costs
- Problems associated with achieving the necessary vacuums
Fig. 5-9 shows the glycol concentrations that can be obtained with various reboiler temperatures.
Fig. 5-11. Shows the effect of stripping gas on TEG concentration
Fig. 5.12. Shows the effect of stripping gas on the concentration using Stahl column
Fig. 5-16. Simplified Process Flow Diagrams of Enhanced TEG Regeneration Systems.
Fig. 5-17. Two different overhead vapor configurations for Stahl column; (a) Vapour flows directly into the bottom of the regeneration column; and (b) Vapor flows into reboiler of the regeneration column.
5.4.2 DRIZO® (wt.-2) Process
DRIZO®, achieves glycol enrichment by means of its own internally generated stripping medium, a mixture of paraffinic and aromatic (BTEX) hydrocarbons of a C5+ boiling range, which are absorbed by the glycol. The heavy hydrocarbons and water are condensed from the regenerator overhead while the non-condensables discharge to atmosphere essentially free of BTEX. The condensed hydrocarbons (with BTEX) are separated from the water, are vaporized and superheated, and flow to the lean-glycol stripping column to serve as the stripping medium, (The mixture of hydrocarbons form an azeotrope with water in the reconcentrator, thus lowering the effective boiling point of the mixture.), which results in TEG purities higher than 99.99 weight percent. As the liquid hydrocarbons build up, they are drawn off as an NGL product. Various options to further enhance the lean TEG purity are available, such as drying the hydrocarbon liquid solvent with solid desiccant, which achieves TEG purities as high as 99.999 weight percent. It is claimed that this high TEG purity level permits water dewpoint depressions in the range of 250°F.
Economic considerations:
- May be favored over stripping gas.
- Existing units can be retrofitted to increase the dehydration capacity.
- Must be evaluated on a case-by-case basis since Drizo (wt.-2) is a Dow-patented process and a license fee is required.
As shown in Figure 5-18, the Drizo process is the same as a conventional TEG dehydration system until the wet glycol flows into the reconcentrator.
Wet glycol is reconcentrated to 98.5% by conventional distillation.
The semi-lean glycol is then counter currently contacted with hydrocarbon solvent (iso-octane) vapors at 400 0F.
The hydrocarbon and water are taken overhead, condensed, and then phase separated.
The water is discarded, and the solvent is recycled into the system.
Drizo system is Competitive with applications that utilize a conventional TEG unit with stripping gas and it is most competitive in the range of -40 0F to -80 0F.
Fig. 5-18 Dow Drizo (wt.-2) gas dehydration process.
Fig. 5-19. Simplified Process Flow Diagrams of Enhanced TEG Regeneration Systems.
5.4.3 Coldfinger® Process
Based on the water TEG, vapor liquid equilibrium diagrams which shows that for any TEG liquid concentration the vapor concentration is richer in water.
Incorporates a closed vessel one-half filled with vapor and liquid at equilibrium with a condenser tube bundle in the vapor space (Figure 5-20).
Fig. 5-20 Cold finger condenser.
The condenser causes water condensation, which is removed from the vessel to a trough placed under the condenser tube bundle. As the condensate is removed:
- System’s equilibrium is upset
- Liquid phase releases more water to the vapor in order to reestablish equilibrium Consequently, the liquid phase has a lower water content than it did originally.
With a limited residence time, the water in the liquid phase is exhausted, and the residual liquid (lean TEG) approaches better than 99.7% by weight TEG concentration. In the most common applications, dilute glycol (rich TEG) from the glycol contactor is used as the coolant in the Coldfinger® tube bundle. The Coldfinger® process does not use stripping gas.
Numerous variations based on this principle exist. One design is shown in Figure 5-21.
Contact between gas and glycol is the same as in a conventional TEG system.
Wet glycol leaves the contractor and flows to the condenser-tube bundle of the cold finger, where it acts as a coolant, and it is used as a coolant in the glycol still before the hydrocarbon liquid phase, hydrocarbon vapor phase, and glycol/water phase are separated in a three-phase separator.
The glycol/water phase is mixed with the cold finger condensate, and is heated by the cold finger liquid product before it is fed to the still. The hot semi-lean glycol (which is near its boiling point) from the still bottoms is fed to the cold finger. The liquid product is cooled, pumped, cooled again, and fed to the contactor.
The main benefit of this system is that it is more fuel-efficient then the conventional TEG system.
However, it is more complex and not as well-proven as the conventional system.
5.5 Other Considerations of Glycol Dehydration
Under conventional dehydration conditions, 40 to 60% of methanol in the feed gas to a glycol dehydrator will be absorbed by the TEG. This will add additional heat duty on the reboiler and additional vapor load on the regenerator. High methanol injection rates and slug carryover can cause flooding.
Glycol losses can be defined as mechanical carryover from the contactor (normally 0.10 gallon/MMscf for standard mist eliminator) plus vaporization from the contactor and regenerator and spillage. Glycol losses, exclusive of spillage, range from 0.05 gallon/MMscf for high pressure low-temperature gases to as much as 0.30 gal/MMscf for low pressure, high temperature gases. Excessive losses usually result from foaming in the absorber and/or regenerator. Anti-foam agents are sometimes used.
TEG vaporization losses at the contactor are minimal unless the gas temperature exceeds about 120°F. These losses are more significant at lower pressures. Tetraethylene glycol (TTEG) has been used in some cases to minimize losses in high temperature, low pressure systems. Vaporization losses at the regenerator typically result from excessive stripping gas rates and/or inadequate reflux.
Glycol losses in CO2 dehydration systems can be significantly higher than in natural gas systems particularly at pressures above about 900 psia. This is due to the solubility of TEG in dense phase CO2. Glycerol is much less soluble and has been used successfully as a desiccant in some CO2 dehydration systems.
Glycol becomes corrosive with prolonged exposure to oxygen. A dry gas blanket on the glycol surge tank will help eliminate oxygen absorption. Special precautions should be taken if oxygen is in the gas to be dehydrated. Thermal decomposition of TEG can become a problem if TEG is heated to temperatures above 400°F.
A low pH accelerates decomposition of glycols. Bases such as triethanolamine, borax, or sodium mercaptobenzothiazole may be added to maintain pH, but they should be added sparingly.
Table 5-2. Glycol regeneration process.
Fig. 5-21. Cold finger condenser process.
5.5.1 Systems Utilizing Glycol-Gas Powered Pumps
Cool wet glycol leaves the bottom of the contactor, passes through a strainer, and powers the pump. The wet glycol takes a pressure drop through the pump, then passes through the reflux condenser coil in the reconcentrator still column.
Fig. 5-22 System utilizing glycol-gas powered pumps.
5.5.2 Systems Utilizing Electric Driven Pumps
Cool wet glycol leaves the bottom of the contactor, passes through a choke and level control operated motor valve, where a pressure drop occurs.
The glycol then passes through the still column reflux coils. From the reflux coil it flows through the first dry glycol to wet glycol heat exchanger, then into the gas/glycol/condensate separator, where insoluble hydrocarbons are removed.
From the separator, the glycol passes through the filter to remove tarry hydrocarbons, then through the second dry glycol to wet glycol heat exchanger and into the reconcentrator still column.
In the top of the reboiler still column, cool wet glycol flows through the reflux condenser coils, preventing glycol from leaving as a vapor. The wet glycol enters the column below the coils and spills downward through the packing and into the reconcentrator.
Heat is circulated through the tube to boil the water from the glycol. A weir holds the level of glycol above the heating tubes. Regenerated glycol flows over the weir and leaves through the outlet in the bottom.
Fig. 5-23 Systems utilizing electric driven pumps.
-------------------Fundamentals of Oil and Gas Processing Book
Basics of Gas Field Processing Book
Prediction and Inhibition of Gas Hydrates Book
Basics of Corrosion in Oil and Gas Industry Book
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Chapter 5 113
Gas Dehydration 113
Introduction 113
Absorption 114
5.1. Gas Dehydration by Absorption 114
5.1.1 Absorption and Stripping 114
5.1.2 Raoult and Dalton’s Laws 114
5.1.3 Glycol-Water Equilibrium 115
5.2 Glycol Dehydration 115
5.2.1 Principles of Operation 115
5.3 Effect of Operating Variables 123
5.3.1 Glycol Selection 123
5.3.2 Inlet Gas Temperature 124
5.3.3 Lean Glycol Temperature 124
5.3.4 Glycol Reconcentrator Temperature 124
5.3.5 Temperature at Top of Still Column 125
5.3.6 Contactor Pressure 125
5.3.7 Reconcentrator Pressure 125
5.3.8 Glycol Concentration 126
5.3.9 Glycol Circulation Rate 128
5.3.10 Number of Absorber Trays 129
5.4 Enhanced Glycol Concentration Processes 132
5.4.2 DRIZO® (wt.-2) Process 133
5.4.3 Coldfinger® Process 135
5.5 Other Considerations of Glycol Dehydration 136
5.5.1 Systems Utilizing Glycol-Gas Powered Pumps 138
5.5.2 Systems Utilizing Electric Driven Pumps 139
5.6 Glycol Gas Contactors 140
5.6.1 Trays and Packing 140
5.7 Glycol Dehydration System Sizing 146
5.7.1 Contactor Diameter 146
5.7.2 Number of Trays and Tray Spacing 148
5.7.3 Downcomers 148
5.7.4 Glycol Circulation Rate 149
5.7.5 Lean Glycol Concentration 149
5.7.6 Glycol-Glycol Preheater 149
5.7.7 Glycol-Gas Cooler 149
5.7.8 Glycol-Glycol Heat Exchanger 149
5.7.9 Gas-Glycol-Condensate Separator 150
5.7.10 Reconcentrator 150
5.7.11 Heat Duty 150
5.7.12 Fire Tube Sizing 151
5.7.13 Reflux Condenser 154
5.7.14 Stripping Still Column 154
5.7.15 Filters 155
5.8 Calculation Examples for Glycol Dehydration 158
Example 5-1 158
Example 5-2 159
Example 5-3 159
5.9 Glycol Unit Operation 161
5.9.1 Start up 161
5.9.2 Routing Operation 161
5.9.3 Shut Down 161
5.10 Glycol Maintenance and Care 162
5.10.1 Preventive Maintenance 162
5.11 Glycol Operation Considerations 165
5.11.1 Oxidation 165
5.11.2 Thermal Decomposition 165
5.11.3 pH Control 166
5.11.4 Salt Contamination and Deposits 166
5.11.5 Hydrocarbons 166
5.11.6 Sludge 167
5.11.7 Foaming and defoamers 167
5.12 Analysis and Control of Glycol 167
5.12.1 Visual Inspection 167
5.12.2 Chemical Analysis 167
5.12.3 Chemical Analysis Interpretation 168
5.13 Troubleshooting 5.13.1 General Considerations 172
5.13.2 Main approach to troubleshooting: 172
5.13.3 High Dew Points 172
5.13.3 Glycol Loss from the Contactor 173
5.13.5 Glycol Loss from the Reconcentrator 174
5.3.16 Glycol Loss From Glycol Hydrocarbon Separator 175
5.13.7 Glycol Loss—Miscellaneous 175
5.14 Glycol System Cleaning 175
5.15 Eliminating Operating Problems 176
5.15.1 Inlet Scrubber/Microfiber Filter Separator 176
5.15.2 Absorber 177
5.15.3 Glycol-Gas Heat Exchanger 178
5.15.4 Lean Glycol Storage Tank or Accumulator 178
5.15.5 Stripper or Still Column 179
5.15.6 Reboiler 180
5.15.7 Stripping Gas 181
5.15.8 Glycol Circulating Pump 182
5.15.9 Flash Tank or Glycol-Gas Separator 183
5.15.10 Gas Blanket 183
5.15.11 Reclaimer 184
5.16 Improving Glycol Filtration 184
Adsorption 185
5.17 Overview of Adsorption Processes 185
5.18 Properties of Industrial Adsorbents for Dehydration 187
5.19 Solid Bed Adsorption Process 188
5.20 Principles of Operation 188
5.20.1 Introduction 188
5.20.2 Drying Cycle 189
5.20.4 Regeneration Cycle 189
5.21 Adsorption System Performance 190
5.22 Effect of Process Variables 191
5.22.1 Quality of Inlet Gas 191
5.22.2 Temperature 191
5.22.3 Pressure 192
5.22.4 Cycle Time 192
5.22.5 Gas Velocities 192
5.22.6 Source of Regeneration Gas 192
5.22.7 Direction of Gas Flow 193
5.22.8 Desiccant Selection 194
5.22.9 Effect of Regeneration Gas on Outlet Gas Quality 195
5.22.10 Pressure Drop Considerations 196
5.22.11 Equipment 197
5.23 Desiccant Performance 200
5.24 Design 202
5.24.1 Pressure Drop & Minimum Diameter 202
5.24.2 Mass desiccant Required & Bed Length 203
5.24.3 Regeneration Calculations 205
5.24.4 Design Example 208
5.25 Nonregenerable Dehydrator 211
5.25.1 Calcium Chloride Dehydrator Unit 211
5.25.2 Principles of Operation 211
5.26 Dehydration by Refrigeration 213
5.27 Dehydration by Membrane Permeation 214
5.28 Other Processes 214
5.29 Comparison of Dehydration Processes 215
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Chapter 5
Gas Dehydration
Introduction
After the liquid (free) water has been removed from the gas stream by separation, 25 to 120 lbs of water per MMscf of gas will remain, depending on the temperature and pressure of the gas.
The warmer the inlet gas and the lower the pressure, the more water vapor the gas stream will contain.
Dehydration is the process used to remove water from natural gas and natural gas liquids (NGLs), and is required to:
1- prevent formation of hydrates and condensation of free water in processing and transportation facilities,
2- meet a water content specification, and
3- prevent corrosion
Techniques for dehydrating natural gas, associated gas condensate and NGLs include:
1- Absorption using liquid desiccants,
2- Adsorption using solid desiccants,
3- Dehydration with CaCl2,
4- Dehydration by refrigeration,
5- Dehydration by membrane permeation,
6- Liquid dehydration by gas stripping, and
7- Liquid dehydration by distillation
Fig.5-1. Dehydration of natural gas.
Absorption
5.1. Gas Dehydration by Absorption
5.1.1 Absorption and Stripping
In the absorption process, a hygroscopic liquid is used to contact wet gas and remove the water vapor. Through absorption, the water in a gas stream is dissolved in a relatively pure liquid solvent stream. Normally, between 20 to 115 lbs of water per MMscf of gas must be removed before the required dew point of the gas is met. The reverse process, in which the water in the solvent is transferred into the gas phase, is known as stripping. The terms regeneration, re-concentration, and reclaiming are also used to describe stripping (or purification) because the solvent is recovered for reuse in the absorption step. Absorption and stripping are frequently used in gas processing, gas sweetening, and glycol dehydration.
5.1.2 Raoult and Dalton’s Laws
Absorption can be qualitatively modeled by using Raoult’s and Dalton’s laws.
For a vapor liquid equilibrium system:
Raoult’s Law state that the partial pressure of a component in a vapor phase that is in equilibrium with a liquid is directly proportional to the mole fraction of the component in the liquid phase. Or, the partial pressure of a component in a vapor phase that is in equilibrium with a liquid equal the vapor pressure of its pure component multiply by mole fraction of the component in liquid phase.
Dalton’s Law states that the partial vapor pressure of a component is equal to the total pressure multiplied by its mole fraction in the gas mixture.
Raoult’s Law expressed in equation form is:
pi = PiXi Eq. 5-1
Dalton’s Law expressed in equation form is:
pi = PYi Eq. 5-2
Where:
pi = Partial vapor pressure of component i
Pi = Vapor pressure of pure component i
Xi = Mole fraction of component i in the liquid
P = Total pressure of the gas mixture
Yi = Mole fraction of component i in the vapor
Combining these laws we have:
PYi = piXi
or
pi/P = Yi/Xi Eq. 5-3
Since the pure-component vapor pressure and the total pressure are not affected by composition.
Equation (5-3) is significant. It states that the ratio of the vapor mole fraction to the liquid mole fraction for any component is independent of the concentrations of that component and the other components present. The ratio Yi /Xi is commonly known as the K-value.
Since the pure component vapor pressure increases with temperature, the K-value increases with increasing temperature and decreases with increasing pressure.
In physical terms this means:
• Transfer from the gas phase to the liquid phase (absorption) is more favorable at lower temperature and high pressures.
• Transfer to the gas phase (stripping) is more favorable at higher temperatures and lower pressures.
5.1.3 Glycol-Water Equilibrium
Absorption processes are dynamic and continuous.
Gas flow cannot be stopped to let the vapor and liquid reach equilibrium. Thus, the system must be designed to approach equilibrium as closely as possible while flow continues. This is accomplished by using a trayed or packed contactor in which the gas and liquid are in counter current flow. The closer to 100% equilibrium that a tray or packed section approaches, the higher the tray or packing efficiency.
For example, A common tray efficiency is 25%, meaning that 25% of the water molecules that would have been transferred under equilibrium conditions were actually transferred.
Wet gas enters the bottom of the column and contacts the rich glycol (high water content) just before the glycol leaves the column. The gas encounters leaner glycol as it works its way up the column, contacting the leanest glycol (lowest water content) just before it leaves the column.
The equilibrium based on Dalton’s and Raoult’s Laws can be rearranged as follows:
Yi = Xi (Pi / P) Eq. 5-4
Since Pi/P is constant for constant temperature, the concentration of the water in the gas must be directly proportional to the concentration in the liquid. However, the liquid concentration is constantly changing as water is absorbed. The counter current flow in the contactor makes it possible for the gas to transfer a significant amount of water to the glycol and still approach equilibrium with the leanest glycol concentration.
5.2 Glycol Dehydration
5.2.1 Principles of Operation
The most common liquid used in absorption type dehydration units is triethylene glycol (TEG)
Liquid desiccant dehydration equipment is simple to operate and maintain. It can easily be automated for unattended operation; for example, glycol dehydration at a remote production well. Liquid desiccants can be used for sour gases, but additional precautions in the design are needed due to the solubility of the acid gases in the desiccant solution. At very high acid gas content and relatively higher pressures the glycols can also be “soluble” in the gas.
Glycols are typically used for applications where dew point depressions of the order of 60° to 120°F are required.
Diethylene glycol (DEG), triethylene glycol (TEG), and tetraethylene glycol (TTEG) are used as liquid desiccants, but TEG is the most common for natural gas dehydration. The schematics in Figures 5-2 and 5-3 show the flow through a typical glycol dehydration system. The glycol dehydration process can be discussed in two parts gas system (Figure 5-2) & glycol system (Figure 5-3)
1. Gas System (figure 5-2)
Wet gas enters the unit through the inlet gas scrubber/microfiber filter separator, usually vertical, to remove liquid and solid impurities. After passing through the microfiber filter separator, the gas enters the glycol gas contactor near the bottom of the vessel.
Even if the dehydrator is near a production separator. The inlet gas scrubber upstream the contactor (or filter) will prevent accidental dumping of large quantities of water (fresh or salty), hydrocarbons, treating chemicals or corrosion inhibitors into the glycol contactor. Even small quantities of these materials can result in excessive glycol losses due to foaming, reduced efficiency, and increased maintenance. Integral separators at the bottom of the contactor are common. Figure 5-4.
Fig. 5-2. Gas system
The inside of the contactor contains either packing or several trays with weirs that maintain a specific level of glycol so that the gas must bubble through the glycol as the gas flows up.Different types of trays will be presented.
As the wet gas passes upward through each succeeding tray, it gives up the water vapor to the glycol and becomes progressively drier. Before leaving the contactor the gas passes through a mist extractor to remove glycol that may be trying to leave the gas.
Dry gas exits the contactor at the top and passes through an external glycol gas heat exchanger where it cools the incoming dry glycol to increase its absorption capacity (Figure 5-2).
Some installations incorporate a glycol knockout drum (centrifugal separator) which recovers any glycol that has escaped with the gas through the mist extractor.
The dry gas then leaves the dehydrator unit.
Fig. 5-4. Integral separators at the bottom of the contactor
2- Glycol System (figure 5-3)
The glycol pump pumps up dry concentrated glycol to contactor, and then passes through the glycol gas heat exchanger before entering the contractor tower.
The glycol gas heat exchanger cools the glycol to near the temperature of the gas before the glycol enters the contactor.
It is important that the glycol be near the gas temperature to prevent gas from exceeding equilibrium temperature, and to prevent foaming. Dry glycol from the glycol gas heat exchanger enters the contactor tower and flows across the top tray.
This is the first contact between the glycol and gas. Glycol flows downward through downcomers in the tower, absorbing more water as it passes across each tray.
The downcomer seals the glycol passage into the tray below, thus preventing gas from short-circuiting past the bubble caps.
As the glycol flows downward through each succeeding tray, it becomes wetter with the water it has absorbed from the gas and collects in the bottom of the contactor saturated with water.
As the gas moves upward through each succeeding tray, it becomes drier.
The wet glycol that has accumulated in the bottom of the contactor passes through a strainer (filter), which removes abrasive particles, before flowing through the power side of the glycol pump (energy exchange pumps), where it furnishes the power to pump the dry glycol into the contactor. Power comes from the increased head caused by the absorbed water contained in the rich glycol.
From the glycol gas contactor the cool wet glycol passes through a coil (reflux condenser) in the top of the reboiler still column. The coil cools the vapors leaving the still column and condenses the glycol vapors to liquid.
The glycol liquid droplets gravitate back down the still column to the re-concentrator. The water remains as a vapor and continues on out the top of the still column. The cooling coil is commonly called the reflux condenser.
The slightly warmed wet glycol leaving the reflux condenser passes through the glycol-glycol preheater. The hot dry glycol from the glycol reconcentrator heats the wet glycol further, and in turn further cools the dry glycol before it goes to the glycol pumps.
After leaving the glycol-glycol preheater, the heated wet glycol is sent to a low-pressure gas-glycol-condensate separator, where most of the entrained gas and liquid hydrocarbons that were picked up by the glycol on its path through the contactor are removed.
The heat provided by the glycol-glycol preheater helps in the separation of hydrocarbons from the wet glycol. The hydrocarbon condensate is separated from the glycol by a three-phase gas-glycol-condensate separator (Fig. 5-5), or a vertical three-phase separator as in fig.5-3.
Fig. 5-5. Gas-glycol-condensate separator
After the gas and condensate has been separated in the gas-glycol-condensate separator, the wet glycol passes through a microfiber filter (Fig. 5-6). These filters are used to remove solids, tarry hydrocarbons, or other impurities.
Fig. 5-6. Microfiber filter.
From the microfiber filter the wet glycol enters a charcoal or carbon filter. Activated carbon granules in this filter absorb liquid-entrained hydrocarbons, well-treating chemicals, compressor oils, and other impurities that may cause foaming.
From the charcoal filter, the wet glycol flows through the dry glycol to the wet glycol heat exchanger. This heat exchanger preheats the wet glycol as much as possible before entering the glycol reconcentrator, thus reducing the heat duty of the glycol reconcentrator.
From the glycol/glycol heat exchanger, the wet glycol enters the still column which sits vertically atop the glycol reconcentrator (Fig. 5-7). The inside of the still column is packed with either ceramic saddles or stainless steel pall rings, which are used to add surface area and distribute heat to the incoming glycol.
Fig. 5-7. Still column
The incoming wet glycol spreads out uniformly and drips down through the packed section.
The vapors traveling upward from the glycol reconcentrator heats the packing. As the glycol travels down through the heated packing, water begins to be driven off as steam. Units utilizing efficient heat exchangers may remove as much as 75 to 80% of the water contained in the glycol in the still column before the glycol reaches the reconcentrator.
As water vapor travels up through the still column and exits from the top, it carries with it trapped glycol vapor. To prevent the loss of glycol vapor, the still column utilizes a “reflux condenser” located on the top of the packed still column.
Glycol vapors escaping the still column with the steam are attracted to the film of condensed liquid (primarily water) covering the coil surface area where they too are condensed. The liquid droplets gravitate back down the still column into the reconcentrator for further treating, thus preventing excessive glycol loss due to vaporization.
On some units, the glycol enters the still column below the packed section of the column.
Vaporization takes place in the reconcentrator. The reflux condenser operates the same in both types of still columns. Condensed liquid from the reflux condenser drops back into the packed section providing a liquid film over the upper portion of the packing. Glycol vapors escaping with steam from the reconcentrator must pass through the packed section.
The watery film covering the packing recaptures the glycol vapor, condensing it into droplets, which wash back into the reconcentrator. Thus, more glycol vapor can be recovered in this configuration than in the previously described still column.
Since vaporization occurs primarily in the reconcentrator, the operating temperature is lower in this type of still column. This translates into:
- Greater reflux condensation
- Requires larger heat duty
From the packed still column, the wet glycol drops downward into the reconcentrator. The glycol is heated to a temperature at which most of the remaining water and some of the glycol are vaporized. A heat source heats the glycol to between 350 0Fand 400 0F, where:
- It removes the remaining water.
- It is below the decomposition point of TEG.
The temperature of the glycol in the reconcentrator is critical and must be controlled at this point.
Sources of heat include:
- Direct fired (natural draft/forced draft)
- Waste heat (exhaust gases from compressors or generators)
- Electric heaters
The heated vapor (both glycol and water) rises upward through the still column.
As the mixture passes the cool reflux condenser coils, the glycol vapors condense and drop back down. The water vapor leaves the top of the still column as steam.
Some of the steam will condense, so a downspout is provided to drain the water off.
A weir maintains a level of glycol over the heat source, which:
- Prevents over heating of the tubes
- Prevents premature tube failure
As the glycol is purified, it spills over the weir into a separate compartment. From the reconcentrator, the dry (lean) glycol flows to the accumulator surge tank, where the glycol pump raises it to contactor pressure to start another cycle.
Figure 5.8 shows a typical, simplified flow sheet for a glycol absorption unit. The wet gas passes through an inlet scrubber to remove solids and free liquids, and then enters the bottom of the glycol contactor. Gas flows upward in the contactor, while lean glycol solution (glycol with little or no water) flows down over the trays. Rich glycol absorbs water and leaves at the bottom of the column while dry gas exits at the top. The rich glycol flows through a heat exchanger at the top of the still where it is heated and provides the coolant for the still condenser. Then the warm solution goes to a flash tank, where dissolved gas is removed. The rich glycol from the flash tank is further heated by heat exchange with the still bottoms, and then becomes the feed to the still. The still produces water at the top and a lean glycol at the bottom, which goes to a surge tank before being returned to the contactor.
Fig. 5-8. Schematic of typical glycol dehydrator unit.
Fig. 5-9. Glycol dehydration system
5.3 Effect of Operating Variables
Several operating and design variables have an important effect on the successful operation of a glycol dehydration system.
5.3.1 Glycol Selection
Glycols are the most commonly used liquid desiccants in the absorption process because they are:
1. Highly hygroscopic (readily absorb and retain water)
2. Stable to heat and chemical decomposition at the temperature and pressures necessary in the process Low vapor pressures, which minimize equilibrium loss of the glycol in the residual natural gas stream and in the regeneration systebm
3. Easily regenerated (water removed) for reuse
4. Noncorrosive and nonfoaming at normal conditions; impurities in the gas stream can change this, but even then inhibitors can help to minimize these problems
5. Readily available at moderate cost
Hygroscopicity of glycols is affected by the concentration (glycol-to-water ratio), that is, increasing as the concentration increases.
Dew point depression obtainable in a gas stream increases as the glycol concentration increases.
Ethylene Glycol (EG)
Ethylene Glycol tends to have high vapor losses to gas when used in a contactor. It is used as a hydrate inhibitor where it can be recovered from the gas by separation at temperatures <50 0F.
Diethylene Glycol (DEG)
DEG (Diethylene) is somewhat cheaper to buy and sometimes is used for this reason. But, by the time it is handled and added to the units there is no real saving. Compared to TEG, DEG has a larger carry-over loss, offers less dewpoint depression and regeneration to high concentrations is more difficult. For these reasons, it is difficult to justify a DEG unit, although a few units are built each year. Diethylene Glycol reconcentrates at temperatures between 3150 and 325 0F, which yields purity of 97.0%. It degrades at 3280F. It cannot achieve the concentration required for most applications.
Triethylene Glycol (TEG)
Triethylene Glycol is most commonly used in glycol dehydration. It reconcentrates at temperatures between 3500 and 400 0F, which yields purity of 98.8%.
It degrades at 404 0F. It tends to experience high vapor losses to gas at temperatures in excess of 1200F. With stripping gas, dew point depressions up to 1500F are possible.
TEG (Triethylene) is preferred for use in dehydration units because:
• It is more easily regenerated due to its high boiling point and other physical properties.
• It has lower vaporization losses than other glycols
• It has lower capital and operating costs than other glycol systems
Tetraethylene Glycol (TTEG)
TTEG (Tetraethylen) is more viscous and more expensive than the other processes. It reconcentrates at temperatures between 4000 and 4300F. It experiences lower vapor losses to gas at high gas contactor temperatures. It degrades at 4600F. The Only real advantage is its lower vapor pressure which reduces absorber carry-over loss. It may be used in those relatively rare cases where glycol dehydration will be employed on a gas whose temperature exceeds about 50°C [122°F].
Table 5-1. Physical properties of glycols and methanol.
5.3.2 Inlet Gas Temperature
At constant pressure, the water content of the inlet gas increases as the temperature increases.
For example, at 1000 psia and 800F, gas holds 34 lb of water/MMscf. While, at 1200F, gas holds 104 lb of water/MMscf.
If the gas is saturated at the higher temperature, the glycol will have to remove about three times as much water to meet the specifications.
Temperatures above 1150F result in high glycol losses, thus requires tetraethylene glycol.
Temperature should not fall below the hydrate formation temperature range (650 to 700F) and always-above 500F.
Temperatures below 500F cause problems due to an increase in glycol viscosity.
Temperatures below 600 to 700F can cause a stable emulsion with liquid hydrocarbons in the gas and cause foaming in the contactor.
An increase in gas temperature increases the gas volume, which in turn increases the diameter of the glycol contactor.
5.3.3 Lean Glycol Temperature
Dry glycol temperatures entering the top tray of the contactor (approach temperature) should be held low (100 to 150F) above the inlet gas temperature.
Equilibrium conditions between the glycol and the water vapor in the gas is affected by temperature. Glycol entering the top tray of the contactor may raise the temperature of the gas surrounding it and prevent the gas giving up its remaining water vapor.
Inlet glycol temperatures greater than 150F above the gas temperature results in high glycol losses to the gas. Drastic temperature differential also has a tendency to emulsify the glycol with any contaminants subsequent glycol loss.
5.3.4 Glycol Reconcentrator Temperature
Reconcentrator temperature controls the concentration of the water in the glycol.
With a constant pressure, the glycol concentration increases with higher reconcentrator temperature. Reconcentrator temperature should be limited to between 3500 and 4000F.
Minimizes degradation of TEG which begins to degrade at 4040F. At a 400°F (204°C), the typical maximum regeneration temperature, TEG yields a lean-glycol concentration of 98.6 wt% at sea level. Higher purity requires reduction of the partial pressure of water in the reboiler vapor space. The most common way to achieve this pressure reduction is to use a stripping gas, or vacuum distillation, which yields lean glycol concentrations of 99.95 wt% and 99.98 wt%, respectively (Detailed process description will be followed in “Enhanced Glycol Concentration Processes” section).
5.3.5 Temperature at Top of Still Column
A high temperature in the top of the still column can increase glycol losses due to excessive vaporization. A reboiler temperature in the range of 3500 to 4000F insures adequate heat transfer to the ceramic packing in the still column.
The still column operates best (allows the steam to escape) when the vapor outlet temperature is between 2150 and 2250F.
When the temperature reaches 2500F and above, glycol vaporization losses increase.
Still top temperature can be lowered by increasing the amount of glycol flowing through the reflux condenser coil. If the temperature in the top of the still column drops too low, (below 2200F) too much water can be condensed and washed back into the reconcentrator, which increases the reconcentrator heat duty. Too much cool glycol circulation in the reflux condenser coil can lower the still top temperature below 2200F, which can cause the excess water to condense. Thus, most reflux condenser coils have a bypass valve, which allows manual or automatic control of the stripping still temperature.
5.3.6 Contactor Pressure
At a constant temperature, the water content of the inlet gas decreases with an increase in pressure. The lower the pressure, the larger the contactor diameter required.
Good dehydration can be achieved at any pressure below 3000 psig as long as the pressure is constant. Optimum dehydration pressure is often in the range of 550 to 1200 psig.
Sizing calculations should always be based on minimum expected operating gas pressure.
Rapid pressure changes translate into rapid velocity changes in the contactor which:
- Breaks the liquid seals between the downcomers and the trays
- Allows the gas to channel up through both the downcomer and bubble caps
- Allows the glycol to be swept out with the gas
5.3.7 Reconcentrator Pressure
Reducing the pressure in the reconcentrator at a constant temperature results in higher glycol purity. Most reconcentrators operate between 0.25 to 0.75 psig of pressure.
On standard atmospheric reconcentrators, pressures in excess of 1 psig results in:
- Glycol loss from the still column
- Reduction of lean glycol concentration
- Reduction in dehydration efficiency
Pressures more than 1 psig are usually associated with excess water in the glycol and create a vapor velocity exiting the still great enough to sweep glycol out.
Fouled still column packing often contributes to high reconcentrator pressure.
Still column should be adequately vented and packing replaced periodically to prevent backpressure on the reconcentrator.
Pressures below atmospheric will increase the lean glycol concentration because the boiling temperature of the rich glycol/water mixture decreases.(will be discussed later)
Reconcentrators are rarely operated in a vacuum due to the added complexity and the fact that air leaks will result in glycol degradation.
If lean glycol concentrations in the range of 99.5% are required consider:
Operating the reconcentrator at a pressure 500 mm Hg absolute (10 psia), or Using stripping gas Fig. 5-9 can be used to estimate the effect of operating in a vacuum on lean glycol concentration.
Fig. 5-9. Glycol purity versus reconcentrator temperature at different levels of vacuum.
5.3.8 Glycol Concentration
The water content of the dehydrated gas depends primarily on the lean glycol concentration.
The higher the concentration of lean glycol entering the contactor, the greater the dew point depression for a given circulation rate and number of trays. Increasing the glycol concentration above a 99% purity can lead to dramatic results on the outlet dew point (Fig. 5-10).
For example, with a 100 0F inlet gas temperature (110 0F top tray temperature), an outlet dew point of
10 0F can be obtained with 99.0% TEG,
-30 0F can be obtained with 99.8% TEG
-40 0F can be obtained with 99.9% TEG
Higher concentrations of TEG can be obtained by:
- Increasing the glycol reconcentration temperature
- Injecting stripping gas into the reconcentrator
- Reducing the operating pressure of the reconcentrator
Fig.5-10. Equilibrium water dew points with various concentrations of TEG.
Reconcentration temperatures for TEG normally run between 380 0F and 400 0F, which results in glycol purities of 98% to 99%. Figures 5-11 and 5-12 illustrate the effect of stripping gas.
If gas is injected directly into the reconcentrator (via a sparger tube), the concentration of TEG increases significantly from 99.1% to near 99.6% as the gas rate is increased from 0 to 4 SCF/gal.
When the Stahl method is used (counter current gas stripping after the reconcentrator – will be illustrated here later), concentrations as high as 99.95% TEG can be attained at a 400 0F reconcentrator temperature.
Fig. 5-11. Effect of stripping gas on TEG concentration.
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Fig. 5.12. Effect of stripping gas on the concentration using Stahl column
5.3.9 Glycol Circulation Rate
When the number of absorber trays and lean glycol concentration are held constant, the dew point depression of a saturated gas is increased as the glycol circulation rate is increased.
The more lean glycol that comes into contact with the gas, the more water vapor is absorbed out of the gas. Whereas the glycol concentration mainly affects the dew point of the dry gas, the glycol rate controls the total amount of water that can be removed. The normal operating level in a standard dehydrator is 3 gallons of glycol per pound of water removed (Range 2-7).
Figure 5-13 shows that a greater dew point depression is a factor of both glycol concentration and glycol circulation rate.
Fig. 5-13. Calculated dew point depression versus circulation rate (1 equilibrium tray (4 actual trays)).
Excessive circulation rates:
- Overload the reconcentrator and/or decrease the reconcentrator temperature
- Prevent good glycol regeneration, resulting in decrease lean glycol concentration.
- Prevent an adequate glycol gas contact in the absorber; this may results in a decrease of water removed by glycol from the gas.
- Increase pump maintenance problems
- Increase glycol loses
- Only if the reconcentrator temperature remains constant, an increase in circulation rate will lower the dew point of the gas.
5.3.10 Number of Absorber Trays
When the glycol circulation rate and the lean glycol concentration are held constant, the dew point depression of a saturated gas is increased as the number of trays is increased.
Actual trays do not reach equilibrium, and their approach to it is expressed as a fraction of a theoretical tray. A tray efficiency of 25% is commonly used for design.
Four actual trays with efficiencies of 25% would accomplish the job of one theoretical tray.
The number of actual trays in a design ranges from 4 to 12.
An approximation of the actual number of valve trays per foot of packing can be obtained from
Figure 5-14.
Fig. 5-14 Trays of packing required for glycol dehydrator.
For high performance units, the specification of more than 4 trays in a new design can achieve fuel savings (for the same dew point depression) due to:
- Lower circulate rate
- Lower reconcentration temperature
- Lower stripping gas rate
Figure 5-15 shows that specifying a few additional trays in the contactor is much more effective than increasing the glycol circulation rate. The additional investment for a taller contactor is often justified by fuel savings.
Evaluation of a TEG system involves first establishing the minimum TEG concentration required to meet the outlet gas water dewpoint specification. Fig. 5-10 shows the water dewpoint of a natural gas stream in equilibrium with a TEG solution at various temperatures and TEG concentrations. Fig. 5-10 can be used to estimate the required TEG concentration for a particular application or the theoretical dewpoint depression for a given TEG concentration and contactor temperature. Actual outlet dewpoints depend on the TEG circulation rate and number of equilibrium stages, but typical approaches to equilibrium are 10-20°F. Equilibrium dewpoints are relatively insensitive to pressure and Fig. 5-10 may be used up to 1500 psia with little error.
Fig. 5-15 Effect of number of absorber trays on dew point depression.
When dehydrating to very low dewpoints, such as those required upstream of a refrigeration process, the TEG concentration must be sufficient to dry the gas to the hydrate dewpoint.
Once the lean TEG concentration has been established, the TEG circulation rate and number of trays (height of packing) must be determined. Most economical designs employ circulation rates of about 2-5 gal. TEG/lb H2O absorbed. The relationship between circulation rate and number of equilibrium stages (N) employs the absorption calculation techniques set out in Engineering Data Book “GPSA 2004” Chapter 19. Calculation results for TEG systems are presented in figures 5-36 through 5-41 where they show the fraction of water removed versus TEG rate with respect to different glycol purities. (N=1 theoretical trays, 4 actual trays)
The graphs in these figures apply only if the feed gas is water saturated. They are based on feed-gas and therefore contactor temperatures of 80°F, but are essentially independent of temperature.
Although the K values in the absorption factors (i.e., L/VK) do increase with temperature, the required TEG rates also increase, and this tends to compensate for the increasing K values and keep the absorption factors fairly constant.
Conversion from equilibrium stages to actual trays can be made assuming an overall tray efficiency of 25-30%. For random and structured packing, Height of Packing Equivalent to a Theoretical Plate (HETP) varies with TEG circulation rate, gas rate, gas density, and packing characteristics but a value of about 60 inches is usually adequate for planning purposes.
When the gas density exceeds about 6 lb/ft3 (generally at very high pressures), the above conversions may not provide sufficient packing height and number of trays. Also when the contactor temperature is less than about 60°F, the increased viscosity of the TEG can reduce mass transfer efficiency, and temperatures in this range should be avoided.
Typical tray spacing in TEG contactors is 24 inches. Therefore, the total height of the contactor column will be based on the number of trays or packing required plus an additional 6-10 ft to allow space for vapor disengagement above the top tray, inlet gas distribution below the bottom tray, and rich glycol surge volume at the bottom of the column. Bubble cap trays have historically been used in glycol contactors due to the low liquid rates versus gas flow, but structured packing is widely used. Generally, structured packing allows a significantly smaller contactor diameter and a slightly smaller contactor height. Calculation examples will be followed.
5.4 Enhanced Glycol Concentration Processes
The enhanced glycol concentration processes include stripping gas, vacuum, Drizo, and cold finger.
5.4.1 Stripping Gas and Vacuum
Normal dehydration systems with TEG glycol purity of 98.5% are capable of achieving dew point depressions up to 70 0F.
If very pure glycol (up to 99.9% TEG) is required and cannot be achieved by the standard regeneration system, stripping gas may be used. A small amount of dry natural gas, normally taken from the fuel stream, is injected into the reconcentrator. Since hot gas has an affinity for water, the stripping gas is bubbled through the hot glycol, which strips the remaining water from the glycol.
This gas can be put directly into the reconcentrator or it can be added to the storage tank where it can percolate through the packed column between the two vessels (Stahl column) fig 5-17.
The Stahl column also serves as a weir where the dry glycol spills downward by gravity over packing while the gas goes upward, removing even more water.
This method prevents air from coming into contact with the dry glycol in the storage tank, thus preventing oxidization of the glycol.
Oxygen entry into the glycol system will:
- Decompose the glycol to some extent
- Cause corrosion within the system.
Stripping gas can:
- Reduce the temperature at which the reconcentrator must operate
- Reduce the glycol circulation rate necessary to dehydrate the gas adequately
Stripping gas may be used to obtain higher dew point depressions.
Vacuum-operated glycol units can achieve glycol purities of up to 99.9% but are rarely used because of:
- High operating costs
- Problems associated with achieving the necessary vacuums
Fig. 5-9 shows the glycol concentrations that can be obtained with various reboiler temperatures.
Fig. 5-11. Shows the effect of stripping gas on TEG concentration
Fig. 5.12. Shows the effect of stripping gas on the concentration using Stahl column
Fig. 5-16. Simplified Process Flow Diagrams of Enhanced TEG Regeneration Systems.
Fig. 5-17. Two different overhead vapor configurations for Stahl column; (a) Vapour flows directly into the bottom of the regeneration column; and (b) Vapor flows into reboiler of the regeneration column.
5.4.2 DRIZO® (wt.-2) Process
DRIZO®, achieves glycol enrichment by means of its own internally generated stripping medium, a mixture of paraffinic and aromatic (BTEX) hydrocarbons of a C5+ boiling range, which are absorbed by the glycol. The heavy hydrocarbons and water are condensed from the regenerator overhead while the non-condensables discharge to atmosphere essentially free of BTEX. The condensed hydrocarbons (with BTEX) are separated from the water, are vaporized and superheated, and flow to the lean-glycol stripping column to serve as the stripping medium, (The mixture of hydrocarbons form an azeotrope with water in the reconcentrator, thus lowering the effective boiling point of the mixture.), which results in TEG purities higher than 99.99 weight percent. As the liquid hydrocarbons build up, they are drawn off as an NGL product. Various options to further enhance the lean TEG purity are available, such as drying the hydrocarbon liquid solvent with solid desiccant, which achieves TEG purities as high as 99.999 weight percent. It is claimed that this high TEG purity level permits water dewpoint depressions in the range of 250°F.
Economic considerations:
- May be favored over stripping gas.
- Existing units can be retrofitted to increase the dehydration capacity.
- Must be evaluated on a case-by-case basis since Drizo (wt.-2) is a Dow-patented process and a license fee is required.
As shown in Figure 5-18, the Drizo process is the same as a conventional TEG dehydration system until the wet glycol flows into the reconcentrator.
Wet glycol is reconcentrated to 98.5% by conventional distillation.
The semi-lean glycol is then counter currently contacted with hydrocarbon solvent (iso-octane) vapors at 400 0F.
The hydrocarbon and water are taken overhead, condensed, and then phase separated.
The water is discarded, and the solvent is recycled into the system.
Drizo system is Competitive with applications that utilize a conventional TEG unit with stripping gas and it is most competitive in the range of -40 0F to -80 0F.
Fig. 5-18 Dow Drizo (wt.-2) gas dehydration process.
Fig. 5-19. Simplified Process Flow Diagrams of Enhanced TEG Regeneration Systems.
5.4.3 Coldfinger® Process
Based on the water TEG, vapor liquid equilibrium diagrams which shows that for any TEG liquid concentration the vapor concentration is richer in water.
Incorporates a closed vessel one-half filled with vapor and liquid at equilibrium with a condenser tube bundle in the vapor space (Figure 5-20).
Fig. 5-20 Cold finger condenser.
The condenser causes water condensation, which is removed from the vessel to a trough placed under the condenser tube bundle. As the condensate is removed:
- System’s equilibrium is upset
- Liquid phase releases more water to the vapor in order to reestablish equilibrium Consequently, the liquid phase has a lower water content than it did originally.
With a limited residence time, the water in the liquid phase is exhausted, and the residual liquid (lean TEG) approaches better than 99.7% by weight TEG concentration. In the most common applications, dilute glycol (rich TEG) from the glycol contactor is used as the coolant in the Coldfinger® tube bundle. The Coldfinger® process does not use stripping gas.
Numerous variations based on this principle exist. One design is shown in Figure 5-21.
Contact between gas and glycol is the same as in a conventional TEG system.
Wet glycol leaves the contractor and flows to the condenser-tube bundle of the cold finger, where it acts as a coolant, and it is used as a coolant in the glycol still before the hydrocarbon liquid phase, hydrocarbon vapor phase, and glycol/water phase are separated in a three-phase separator.
The glycol/water phase is mixed with the cold finger condensate, and is heated by the cold finger liquid product before it is fed to the still. The hot semi-lean glycol (which is near its boiling point) from the still bottoms is fed to the cold finger. The liquid product is cooled, pumped, cooled again, and fed to the contactor.
The main benefit of this system is that it is more fuel-efficient then the conventional TEG system.
However, it is more complex and not as well-proven as the conventional system.
5.5 Other Considerations of Glycol Dehydration
Under conventional dehydration conditions, 40 to 60% of methanol in the feed gas to a glycol dehydrator will be absorbed by the TEG. This will add additional heat duty on the reboiler and additional vapor load on the regenerator. High methanol injection rates and slug carryover can cause flooding.
Glycol losses can be defined as mechanical carryover from the contactor (normally 0.10 gallon/MMscf for standard mist eliminator) plus vaporization from the contactor and regenerator and spillage. Glycol losses, exclusive of spillage, range from 0.05 gallon/MMscf for high pressure low-temperature gases to as much as 0.30 gal/MMscf for low pressure, high temperature gases. Excessive losses usually result from foaming in the absorber and/or regenerator. Anti-foam agents are sometimes used.
TEG vaporization losses at the contactor are minimal unless the gas temperature exceeds about 120°F. These losses are more significant at lower pressures. Tetraethylene glycol (TTEG) has been used in some cases to minimize losses in high temperature, low pressure systems. Vaporization losses at the regenerator typically result from excessive stripping gas rates and/or inadequate reflux.
Glycol losses in CO2 dehydration systems can be significantly higher than in natural gas systems particularly at pressures above about 900 psia. This is due to the solubility of TEG in dense phase CO2. Glycerol is much less soluble and has been used successfully as a desiccant in some CO2 dehydration systems.
Glycol becomes corrosive with prolonged exposure to oxygen. A dry gas blanket on the glycol surge tank will help eliminate oxygen absorption. Special precautions should be taken if oxygen is in the gas to be dehydrated. Thermal decomposition of TEG can become a problem if TEG is heated to temperatures above 400°F.
A low pH accelerates decomposition of glycols. Bases such as triethanolamine, borax, or sodium mercaptobenzothiazole may be added to maintain pH, but they should be added sparingly.
Table 5-2. Glycol regeneration process.
Fig. 5-21. Cold finger condenser process.
5.5.1 Systems Utilizing Glycol-Gas Powered Pumps
Cool wet glycol leaves the bottom of the contactor, passes through a strainer, and powers the pump. The wet glycol takes a pressure drop through the pump, then passes through the reflux condenser coil in the reconcentrator still column.
Fig. 5-22 System utilizing glycol-gas powered pumps.
5.5.2 Systems Utilizing Electric Driven Pumps
Cool wet glycol leaves the bottom of the contactor, passes through a choke and level control operated motor valve, where a pressure drop occurs.
The glycol then passes through the still column reflux coils. From the reflux coil it flows through the first dry glycol to wet glycol heat exchanger, then into the gas/glycol/condensate separator, where insoluble hydrocarbons are removed.
From the separator, the glycol passes through the filter to remove tarry hydrocarbons, then through the second dry glycol to wet glycol heat exchanger and into the reconcentrator still column.
In the top of the reboiler still column, cool wet glycol flows through the reflux condenser coils, preventing glycol from leaving as a vapor. The wet glycol enters the column below the coils and spills downward through the packing and into the reconcentrator.
Heat is circulated through the tube to boil the water from the glycol. A weir holds the level of glycol above the heating tubes. Regenerated glycol flows over the weir and leaves through the outlet in the bottom.
Fig. 5-23 Systems utilizing electric driven pumps.
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Re: Basics of Gas Field Processing Book "Full text"
Chapter 5- Part 1 B
5.6 Glycol Gas Contactors
There are two basic types of glycol gas contactor:
1- Trayed towers
2- Packed towers
Some contactors have an “internal scrubber” which occupies approximately the lower one-third of the vessel (Figure 5-4). They are usually installed on units where the inlet gas flow rate is less than 50
MMscfd. “Chimney” is included on the scrubber/contactor combination:
Consists of a large stack that covers the top of the inlet scrubber.
Allows the gas to pass upward from the scrubber section to the absorber section. Prevents glycol from being lost out of the scrubber section.
Some contact towers have an internal three phase separator:
Distinguishable in that the lower section has two sets of level controls and two liquid dump valves, this design is not recommended as it is difficult to operate and maintain.
A separate two-phase microfiber filter separator located immediately upstream of the contactor is the most efficient configuration
The contactor is usually a tray column containing 4 to 12 trays on which up flowing gas bubbles through down flowing glycol, The number of trays in the contactor will affect the amount of moisture removed from the gas by the glycol; more trays mean more moisture removal.
In smaller capacity units, that is, contactors having a diameter of 18 inches or less, random packing may be used instead of trays. The packing is metal, plastic, or ceramic structures that are designed to furnish a large surface area for the glycol solution to spread out and make better contact with the gas. Random packing is poured into the contactor onto a support gad.
Four feet packing is usually standard and sufficient for dew point depressions up to 55°F to 65°F (13°C to 18°C). If higher dew point depressions are required, additional packing may be required.
Packed columns utilize the same process as tray columns, that is, glycol flows down over the packing and gas flows up through the packing contacting the glycol. Packed columns are less expensive; however, the glycol tends to channel easier and have poorer flow distribution. Therefore, special attention must be given to the design of a glycol distribution header above the packing so that gas/glycol contacting will be continuous throughout the packing and the glycol will not channel.
Contactors being designed today may contain structured packing. Structured packing is a group of corrugated metal sheets welded into a specific pattern and placed in the contactor on edge. Glycol coats these sheets and the gas flows between them. This type of packing is much more efficient than bubble caps or random packing. Structured packing is used in columns from six inches (15 centimeters) in diameter up to ten feet (three meters).
5.6.1 Trays and Packing
The more stages, the more complete the absorption, but the taller and more costly the tower.
5.6.1.1 Trays
For most trays, liquid flows across an “active area” of the tray and then into a “down-comer” to the next tray below, etc. Inlet and/or outlet weirs control the liquid distribution across the tray. Vapor flows up the stabilizer tower and passes through the tray active area, bubbling up through (and thus contacting) the liquid flowing across the tray. The vapor distribution is controlled by:
• Perforations in the tray deck (sieve trays),
• Bubble caps (bubble cap trays), or
• Valves (valve trays).
Sieve Trays
Sieve trays are the least expensive tray option. In sieve trays, vapor flowing up through the tower contacts the liquid by passing through small perforations in the tray floor (Figure 5-24). Sieve trays rely on vapor velocity to exclude liquid from falling through the perforations in the tray floor. If the vapor velocity is much lower than design, liquid will begin to flow through the perforations rather than into the downcomer.
This condition is known as weeping. Where weeping is severe, the equilibrium efficiency will be very low. For this reason, sieve trays have a very small turndown ratio.
Fig. 5-24. Vapor flow through a sieve tray.
Valve Trays
Valve trays are essentially modified sieve trays. Like sieve trays, holes are punched in the tray floor. However, these holes are much larger than those in sieve trays. Each of these holes is fitted with a device called a “valve.” Vapor flowing up through the tower contacts the liquid by passing through valves in the tray floor (Figure 5-25). Valves can be fixed or moving. Fixed valves are permanently open and operate as deflector plates for the vapor coming up through the tray floor. For moving valves, vapor passing through the tray floor lifts the valves and contacts the liquid. Moving valves come in a variety of designs, depending on the manufacturer and the application. At low vapor rates, valves will close, helping to keep liquid from falling through the holes in the deck.
At sufficiently low vapor rates, a valve tray will begin to weep. That is, some liquid will leak through the valves rather than flowing to the tray down-comers. At very low vapor rates, it is possible that all the liquid will fall through the valves and no liquid will reach the down-comers. This severe weeping is known as “dumping.” At this point, the efficiency of the tray is nearly zero.
Fig. 5-25. Vapor flow through valve tray
Bubble Cap Trays
In bubble cap trays, vapor flowing up through the tower contacts the liquid by passing through bubble caps (Figure 5-26).
Each bubble cap assembly consists of a riser and a cap. The vapor rising through the tower passes up through the riser in the tray floor and then is turned downward to bubble into the liquid surrounding the cap. Because of their design, bubble cap trays cannot weep. However, bubble cap trays are also more expensive and have a lower vapor capacity/higher pressure drop than valve trays or sieve trays.
Figure 5-26. Vapor flow through bubble cap tray
Tray Design
Bubble Cap Trays (Figures 5-26 through 5-29) Most commonly used design
Better than conventional packing (Figures 5-30 and 5-31)
Fig. 5-27. Bubble cap components.
Fig. 5-28. Bubble cap tray.
Fig. 5-29 Bubble cap tray inside the contactor tower.
Bubble Cap Trays vs. Valve Trays
At low vapor rates, valve trays will weep. Bubble cap trays cannot weep (unless they are damaged). For this reason, it is generally assumed that bubble cap trays have nearly an infinite turndown ratio. This is true in absorption processes (e.g., glycol dehydration), in which it is more important to contact the vapor with liquid than the liquid with vapor. However, this is not true of distillation processes (e.g., stabilization), in which it is more important to contact the liquid with the vapor.
As vapor rates decrease, the tray activity also decreases. There eventually comes a point at which some of the active devices (valves or bubble caps) become inactive. Liquid passing these inactive devices gets very little contact with vapor. At this point, it is possible that liquid may flow across the entire active area without ever contacting a significant amount of vapor. This will result in very low efficiencies for a distillation process. Nothing can be done with a bubble cap tray to compensate for this.
However, a valve tray can be designed with heavy valves and light valves. At high vapor rates, all the valves will be open. As the vapor rate decreases, the valves will begin to close. With light and heavy valves on the tray, the heavy valves will close first, and some or all of the light valves will remain open. If the light valves are properly distributed over the active area, even though the tray activity is diminished at low vapor rates, what activity remains will be distributed across the tray. All liquid flowing across the tray will contact some vapor, and mass transfer will continue. Of course, even with weighted valves, if the vapor rate is reduced enough, the tray will weep and eventually become inoperable.
However, with a properly designed valve tray this point may be reached after the loss in efficiency of a comparable bubble cap tray. So, in distillation applications, valve trays can have a greater vapor turndown ratio than bubble cap trays.
5.6.1.2 Packing
Packing typically comes in two types: random and structured. Liquid distribution in a packed bed is a function of the internal vapor/liquid traffic, the type of packing employed, and the quality of the liquid distributors mounted above the packed bed. Packing material can be plastic, metal, or ceramic. Packing efficiencies can be expressed as height equivalent to a theoretical plate (HETP).
Random Packing
A bed of random packing typically consists of a bed support (typically a gas injection support plate) upon which pieces of packing material are randomly arranged (they are usually poured or dumped onto this support plate). Bed limiters, or hold-downs, are sometimes set above random beds to prevent the pieces of packing from migrating or entraining upward. Random packing comes in a variety of shapes and sizes. For a given shape (design) of packing, small sizes have higher efficiencies and lower capacities than large sizes.
Figure 5-30 shows a variety of random packing designs. An early design is known as a Rasching ring. Rasching rings are short sections of tubing and are low-capacity, low-efficiency, high-pressure drop devices. Today’s industry standard is the slotted metal (Pall) ring. A packed bed made of 1-inch slotted metal rings will have a higher mass transfer efficiency and a higher capacity than will a bed of 1-inch Rasching rings. The HETP for a 2-inch slotted metal ring in a stabilizer is about 36 inches. This is slightly more than a typical tray design, which would require 34 inches (1.4 trays × 24-inch tray spacing) for one theoretical plate or stage.
Structured Packing
A bed of structured packing consists of a bed support upon which elements of structured packing are placed. Beds of structured packing typically have lower pressure drops than beds of random packing of comparable mass transfer efficiency. Structured packing elements are composed of grids (metal or plastic) or woven (metal or plastic) or of thin vertical crimped sheets (metal, plastic, or ceramic) stacked parallel to each other. Figure 5-31 shows examples of the vertical crimped sheet style of structured packing. The grid types of structured packing have very high capacities and very low efficiencies, and are typically used for heat transfer or for vapor scrubbing. The wire mesh and the crimped sheet types of structured packing typically have lower capacities and higher efficiencies than the grid type.
Trays or Packing ?
There is no umbrella answer. The choice is dictated by project scope (new tower or retrofit), current economics, operating pressures, anticipated operating flexibility, and physical properties.
Distillation Service
For distillation services, as in hydrocarbon stabilization, tray design is well understood, and many engineers are more comfortable with trays than with packing. In the past, bubble cap trays were the standard. However, they are not commonly used in this service anymore. Sieve trays are inexpensive but offer a very narrow operating range when compared with valve trays. Although valve trays offer wider operating range than sieve trays, they have moving parts and so may require more maintenance. High capacity/high efficiency trays can be more expensive than standard valve trays. However, high capacity/high efficiency trays require smaller diameter stabilization towers, so they can offer significant savings in the overall cost of the distillation tower. Random packing has traditionally been used in small diameter (<20 inches) towers. This is because it is easier and less expensive to pack these small diameter towers. However, random packed beds are prone to channeling and have poor turndown characteristics when compared with trays. For these reasons, trays were preferred for tower diameters greater than 20 inches.
Stripping Service
For stripping service, as in a glycol or amine contactor, bubble cap trays are the most common. In recent years, there has been a growing movement toward crimped sheet structured packing. Improved vapor and liquid distributor design in conjunction with structured packing can lead to smaller-diameter and shorter stripping towers than can be obtained with trays.
Figure 5-30. Various types of random packing.
Figure 5-31. Structured packing can offer better mass transfer than trays.
5.7 Glycol Dehydration System Sizing
Sizing Involve:
Glycol gas contactor diameter
Number of absorber trays (which establishes the tower’s overall height), tray spacing, and downcomers sizing
Glycol circulation rate
Lean glycol concentration
Reconcentrator heat duty
The number of absorber trays, glycol circulation rate, and lean glycol concentration are all inter-related.
5.7.1 Contactor Diameter
Method #1
The minimum diameter for trayed towers and conventional packing can be determined from the following equation:
d2 = 5040 [(T0ZQg)/P] [(ρg / ρL – ρg)(CD/dm)]0.5 Eq. 5-5
Where:
d = Contactor inside diameter, inches
dm = Drop size, microns =120 to 150 micron range
To = Contactor operating temperature, 0R
Qg = Design gas flow rate, MMscfd
P = Contactor operating pressure, psia
CD = Drag coefficient
ρg = Gas density, lb/ft.3 = 2.7 (SP/TZ) or = ρg= 0.093 ((MW)P)/TZ lb/ft3 (Eq. 1-19)
ρL = Glycol density, lb/ft.3 = 70 lb/ft3
Z = Compressibility factor
S = Gas specific gravity (air = 1)
Structured packing can handle higher gas flow rates for the same diameter contactor.
Method #2
Contactor diameter is set by the gas velocity. Sizing is calculated using recommended values for K-factors and C-factors are shown in table. 5-3, and equations 5-6
Table. 5-3. Recommended sizing parameters for TEG contactors.
G = C [ρV (ρL - ρV)]0.5 Eq. 5-6
Where
G = mass velocity, lb /(ft2 • hr)
C = constant from table 5-3,
ρV = Vapor density, lb/ft3
ρL = Liquid density, lb/ft3
where
2 theoretical stages ≅ 8 bubble cap trays @ 24 inch tray Spacing Eq. 5-7
2 theoretical stages ≅10 ft of structured packing Eq. 5-8
Note: Structured packing vendors frequently quote an Fs value for sizing glycol contactors, where Fs is defined in Eq 5-9.
Fs = v (ρv)0.5 Eq. 5-9
Where
v = vapor velocity, ft/sec
ρv= density, lb/ft3
Values of Fs = 2.5 to 3.0 will generally provide a good estimate of contactor diameter for structured packing.
Method #3
Figures 5-32 through 5-35 are correlations prepared by vessel manufacturers that allow graphical solutions of glycol gas contactor diameters.
Fig. 5-32 Determination of contactor diameter—Sivills.
Fig. 5-33 Determination of contactor diameter—Smith Industries.
Fig. 5-34 Determination of contactor diameter—NATCO.
Fig. 5-35 Determination of contactor diameter—BS&B.
5.7.2 Number of Trays and Tray Spacing
6 to 8 trays are used to meet normal dew-point depressions.
12 trays are typically required for high dew-point depressions.
Spacing ranges from 20 to 30 inches
24 inches is preferred, while 30-inch spacing is recommended if foaming is anticipated
5.7.3 Downcomers
Sized for a maximum velocity of 0.25 ft./sec.
5.7.4 Glycol Circulation Rate
For a given dew-point depression, the circulation rate is dependent upon:
Lean glycol concentration
Number of trays
When the lean glycol concentration and number of trays are held constant, the required glycol circulation rate can be determined from the following equation:
L = (ΔW/Wi) Wi Qg/24 Eq. 5-10
Where:
L = Glycol circulation rates, gal/hr
ΔW/ Wi = Circulation ratio, gal TEG/lb H2O (see Figures 5-36, through 5-41)
Wi = Water content of inlet gas, lb H2O/MMscf
W0 = Desired outlet water content, lbH2O/MMscf
ΔW = Wg / W0
Qg = Gas flow rate, MMscfd
5.7.5 Lean Glycol Concentration
Equilibrium water dew points for various concentrations of TEG are shown in Figure 5-10.
Glycol purity (lean glycol concentration) is a function of the temperature of the reconcentrator (Fig. 5-9).
Glycol purity can be increased by:
a- Adding stripping gas
b- Reducing the pressure in the reconcentrator
c- Reducing the glycol circulation rate
5.7.6 Glycol-Glycol Preheater
Cool wet glycol from the contactor enters the preheater (heat exchanger) at 100 0F and the warm glycol leaves at 1750 to 200 0F en route to the gas/glycol/condensate separator.
Hot dry glycol from the glycol/glycol heat exchanger enters the preheater at 250 0F and the warm dry glycol leaves at 150 0F to the glycol pumps en route to the contactor.
Temperature limitations to the glycol pump:
Glycol powered pumps (Kimray) limited to 200 0F.
Electric plunger pumps limited to 250 0F.
Overall heat transfer coefficient (U = 10 to 12)
5.7.7 Glycol-Gas Cooler
TEG to gas contactor is limited to 10 0F to 15 0F above the inlet gas temperature. If hotter, some TEG will vaporize with gas. If colder, gas condensation of the hydrocarbons may cause foam and glycol loss.
Overall heat transfer coefficient (U = 45).
5.7.8 Glycol-Glycol Heat Exchanger
Hot dry glycol from the reconcentrator enters the heat exchanger at 390 0F and leaves at 250 0F en route to the glycol/glycol preheater. Warm wet glycol from the charcoal filter enters the heat exchanger at 200 0F and the hot wet glycol leaves at 350 0F en route to the still column.
5.7.9 Gas-Glycol-Condensate Separator
Separator should be sized using procedures for sizing gas-liquid separation.
Liquid retention times between 20 and 30 minutes are recommended, depending on API gravity of the condensate. Operating pressure of 35 to 50 psig is recommended.
5.7.10 Reconcentrator
The reconcentrator should be designed to operate 350 0 to 400 0F with TEG, and 305 0F with DEG. Design temperature should be sufficiently below the decomposition point so that hot spots on the fire tube and poor mixing in the reconcentrator will not cause decomposition of the glycol. With everything else operating normally, the reconcentrator temperature is raised to lower the water content of the treated gas, and vice versa. Specific reconcentrator operating temperature is determined by trial and error.
Temperatures up to 400 0F are common
400 0F yields 99.5% TEG purity
375 0F yields 98.3% TEG purity
5.7.11 Heat Duty
Estimated from the following equations
qt = LQL Eq. 5-11
Where:
qt = Total heat duty on reconcentrator, Btu/hr
L = Glycol circulation rate, gal/hr
QL = Reconcentrator heat load, Btu/gal TEG (Table 5-4)
Heat duty estimated from Equation (5-11 above) is normally increased by 10 to 25% to account for start-up, fouling, and increased circulation.
Table. 5-4. Reconcentrator Heat Load
Alternatively, Heat duty can be determined as follows:
Total heat duty =
Sensible Heat, Qs + water vaporization heat duty, QV+ Condenser Duty, QC, Eq. 5-12
Qs = m Cp Δt Eq. 5-13
Where
Qs = Sensible Heat, Btu / gal
m = density lb/gal (table 5-1)
CP = Heat capacity, (Btu/lb.°F)
CP (95.1% TEG) = 0.56 at 110 0F (from physical property of TEG), 0.63 Btu/hr 0F at 200 0F, and 0.70 at 300 0F.- (use 0.67 as an average for the range of 200-300 0F).
Δt = Temperature difference (0F)
Vaporization of Absorbed H2O:
Qv = (ΔHvap) (ΔW) Eq. 5-14
Where
Qv = vaporization of water heat duty, Btu/gal.
ΔHvap = latent heat of vaporization, Btu/lb
ΔW =change of water content
Condenser Duty @ 25% Reflux Ratio:
Qc = % Reflux Ratio X Qv /100 Eq. 5-15
5.7.12 Fire Tube Sizing
The actual surface area of the firetube required for direct-fired heaters can be calculated from the following equation:
A = qt / 6000 Eq. 5-16
Where:
A = Total firetube surface area, ft.2
qt = Total heat duty on reconcentrator, Btu/hr
By determining the diameter and overall length of the U-tube fire tube, one can estimate the overall size of the reconcentrator. A heat flux of 6000 to 8000 Btu/hr-ft.2 is often used, but the 6000 value is suggested to ensure against glycol decomposition.
Fig. 5-36. Water Removal vs. TEG Circulation Rate at Various TEG Concentrations (N = 1.0)
Fig. 5-37. Water Removal vs. TEG Circulation Rate at Various TEG Concentrations (N = 1.5)
Fig. 5-38 . Water Removal vs. TEG Circulation Rate at Various TEG Concentrations (N = 2.0)
Fig. 5-39 . Water Removal vs. TEG Circulation Rate at Various TEG Concentrations (N = 2.5)
Fig. 5-40. Water Removal vs. TEG Circulation Rate at Various TEG Concentrations (N = 3.0)
Fig. 5-41. Water Removal vs. TEG Circulation Rate at Various TEG Concentrations (N = 4.0)
5.7.13 Reflux Condenser
Wet glycol inlet from the gas contactor enters at 115 0F and leaves at 125 0F.
It controls TEG losses. Reflux rate should be 50% of the water removal rate. The condensing coil provides lowest TEG loss and most economical reconcentrator operation.
5.7.14 Stripping Still Column
Temperature considerations:
Temperature is critical to the operation of the still column. Heat is provided by the reconcentration. Reconcentrator temperatures in the range of 350 0F to 400 0F insures adequate heat transfer to the ceramic packing in the still column.
Still columns whose wet glycol inlet enters above the packed section (Figure 5-42):
Operate best with a vapor outlet temperature between 2250 and 250 0F
Purpose of glycol falling over ceramic packing is the efficient use of available heat
Backpressure should be kept to a minimum (1 psig is maximum)
Still columns whose wet glycol inlet enters below the packed section (Figure 5-43):
Allow pall ring type packing to be solely involved in the reflux process
Operate best with a vapor outlet temperature between 185 0F and 195 0F
This temperature allows a greater volume of condensation by the reflux coil while still permitting the majority of the steam to escape
5.6 Glycol Gas Contactors
There are two basic types of glycol gas contactor:
1- Trayed towers
2- Packed towers
Some contactors have an “internal scrubber” which occupies approximately the lower one-third of the vessel (Figure 5-4). They are usually installed on units where the inlet gas flow rate is less than 50
MMscfd. “Chimney” is included on the scrubber/contactor combination:
Consists of a large stack that covers the top of the inlet scrubber.
Allows the gas to pass upward from the scrubber section to the absorber section. Prevents glycol from being lost out of the scrubber section.
Some contact towers have an internal three phase separator:
Distinguishable in that the lower section has two sets of level controls and two liquid dump valves, this design is not recommended as it is difficult to operate and maintain.
A separate two-phase microfiber filter separator located immediately upstream of the contactor is the most efficient configuration
The contactor is usually a tray column containing 4 to 12 trays on which up flowing gas bubbles through down flowing glycol, The number of trays in the contactor will affect the amount of moisture removed from the gas by the glycol; more trays mean more moisture removal.
In smaller capacity units, that is, contactors having a diameter of 18 inches or less, random packing may be used instead of trays. The packing is metal, plastic, or ceramic structures that are designed to furnish a large surface area for the glycol solution to spread out and make better contact with the gas. Random packing is poured into the contactor onto a support gad.
Four feet packing is usually standard and sufficient for dew point depressions up to 55°F to 65°F (13°C to 18°C). If higher dew point depressions are required, additional packing may be required.
Packed columns utilize the same process as tray columns, that is, glycol flows down over the packing and gas flows up through the packing contacting the glycol. Packed columns are less expensive; however, the glycol tends to channel easier and have poorer flow distribution. Therefore, special attention must be given to the design of a glycol distribution header above the packing so that gas/glycol contacting will be continuous throughout the packing and the glycol will not channel.
Contactors being designed today may contain structured packing. Structured packing is a group of corrugated metal sheets welded into a specific pattern and placed in the contactor on edge. Glycol coats these sheets and the gas flows between them. This type of packing is much more efficient than bubble caps or random packing. Structured packing is used in columns from six inches (15 centimeters) in diameter up to ten feet (three meters).
5.6.1 Trays and Packing
The more stages, the more complete the absorption, but the taller and more costly the tower.
5.6.1.1 Trays
For most trays, liquid flows across an “active area” of the tray and then into a “down-comer” to the next tray below, etc. Inlet and/or outlet weirs control the liquid distribution across the tray. Vapor flows up the stabilizer tower and passes through the tray active area, bubbling up through (and thus contacting) the liquid flowing across the tray. The vapor distribution is controlled by:
• Perforations in the tray deck (sieve trays),
• Bubble caps (bubble cap trays), or
• Valves (valve trays).
Sieve Trays
Sieve trays are the least expensive tray option. In sieve trays, vapor flowing up through the tower contacts the liquid by passing through small perforations in the tray floor (Figure 5-24). Sieve trays rely on vapor velocity to exclude liquid from falling through the perforations in the tray floor. If the vapor velocity is much lower than design, liquid will begin to flow through the perforations rather than into the downcomer.
This condition is known as weeping. Where weeping is severe, the equilibrium efficiency will be very low. For this reason, sieve trays have a very small turndown ratio.
Fig. 5-24. Vapor flow through a sieve tray.
Valve Trays
Valve trays are essentially modified sieve trays. Like sieve trays, holes are punched in the tray floor. However, these holes are much larger than those in sieve trays. Each of these holes is fitted with a device called a “valve.” Vapor flowing up through the tower contacts the liquid by passing through valves in the tray floor (Figure 5-25). Valves can be fixed or moving. Fixed valves are permanently open and operate as deflector plates for the vapor coming up through the tray floor. For moving valves, vapor passing through the tray floor lifts the valves and contacts the liquid. Moving valves come in a variety of designs, depending on the manufacturer and the application. At low vapor rates, valves will close, helping to keep liquid from falling through the holes in the deck.
At sufficiently low vapor rates, a valve tray will begin to weep. That is, some liquid will leak through the valves rather than flowing to the tray down-comers. At very low vapor rates, it is possible that all the liquid will fall through the valves and no liquid will reach the down-comers. This severe weeping is known as “dumping.” At this point, the efficiency of the tray is nearly zero.
Fig. 5-25. Vapor flow through valve tray
Bubble Cap Trays
In bubble cap trays, vapor flowing up through the tower contacts the liquid by passing through bubble caps (Figure 5-26).
Each bubble cap assembly consists of a riser and a cap. The vapor rising through the tower passes up through the riser in the tray floor and then is turned downward to bubble into the liquid surrounding the cap. Because of their design, bubble cap trays cannot weep. However, bubble cap trays are also more expensive and have a lower vapor capacity/higher pressure drop than valve trays or sieve trays.
Figure 5-26. Vapor flow through bubble cap tray
Tray Design
Bubble Cap Trays (Figures 5-26 through 5-29) Most commonly used design
Better than conventional packing (Figures 5-30 and 5-31)
Fig. 5-27. Bubble cap components.
Fig. 5-28. Bubble cap tray.
Fig. 5-29 Bubble cap tray inside the contactor tower.
Bubble Cap Trays vs. Valve Trays
At low vapor rates, valve trays will weep. Bubble cap trays cannot weep (unless they are damaged). For this reason, it is generally assumed that bubble cap trays have nearly an infinite turndown ratio. This is true in absorption processes (e.g., glycol dehydration), in which it is more important to contact the vapor with liquid than the liquid with vapor. However, this is not true of distillation processes (e.g., stabilization), in which it is more important to contact the liquid with the vapor.
As vapor rates decrease, the tray activity also decreases. There eventually comes a point at which some of the active devices (valves or bubble caps) become inactive. Liquid passing these inactive devices gets very little contact with vapor. At this point, it is possible that liquid may flow across the entire active area without ever contacting a significant amount of vapor. This will result in very low efficiencies for a distillation process. Nothing can be done with a bubble cap tray to compensate for this.
However, a valve tray can be designed with heavy valves and light valves. At high vapor rates, all the valves will be open. As the vapor rate decreases, the valves will begin to close. With light and heavy valves on the tray, the heavy valves will close first, and some or all of the light valves will remain open. If the light valves are properly distributed over the active area, even though the tray activity is diminished at low vapor rates, what activity remains will be distributed across the tray. All liquid flowing across the tray will contact some vapor, and mass transfer will continue. Of course, even with weighted valves, if the vapor rate is reduced enough, the tray will weep and eventually become inoperable.
However, with a properly designed valve tray this point may be reached after the loss in efficiency of a comparable bubble cap tray. So, in distillation applications, valve trays can have a greater vapor turndown ratio than bubble cap trays.
5.6.1.2 Packing
Packing typically comes in two types: random and structured. Liquid distribution in a packed bed is a function of the internal vapor/liquid traffic, the type of packing employed, and the quality of the liquid distributors mounted above the packed bed. Packing material can be plastic, metal, or ceramic. Packing efficiencies can be expressed as height equivalent to a theoretical plate (HETP).
Random Packing
A bed of random packing typically consists of a bed support (typically a gas injection support plate) upon which pieces of packing material are randomly arranged (they are usually poured or dumped onto this support plate). Bed limiters, or hold-downs, are sometimes set above random beds to prevent the pieces of packing from migrating or entraining upward. Random packing comes in a variety of shapes and sizes. For a given shape (design) of packing, small sizes have higher efficiencies and lower capacities than large sizes.
Figure 5-30 shows a variety of random packing designs. An early design is known as a Rasching ring. Rasching rings are short sections of tubing and are low-capacity, low-efficiency, high-pressure drop devices. Today’s industry standard is the slotted metal (Pall) ring. A packed bed made of 1-inch slotted metal rings will have a higher mass transfer efficiency and a higher capacity than will a bed of 1-inch Rasching rings. The HETP for a 2-inch slotted metal ring in a stabilizer is about 36 inches. This is slightly more than a typical tray design, which would require 34 inches (1.4 trays × 24-inch tray spacing) for one theoretical plate or stage.
Structured Packing
A bed of structured packing consists of a bed support upon which elements of structured packing are placed. Beds of structured packing typically have lower pressure drops than beds of random packing of comparable mass transfer efficiency. Structured packing elements are composed of grids (metal or plastic) or woven (metal or plastic) or of thin vertical crimped sheets (metal, plastic, or ceramic) stacked parallel to each other. Figure 5-31 shows examples of the vertical crimped sheet style of structured packing. The grid types of structured packing have very high capacities and very low efficiencies, and are typically used for heat transfer or for vapor scrubbing. The wire mesh and the crimped sheet types of structured packing typically have lower capacities and higher efficiencies than the grid type.
Trays or Packing ?
There is no umbrella answer. The choice is dictated by project scope (new tower or retrofit), current economics, operating pressures, anticipated operating flexibility, and physical properties.
Distillation Service
For distillation services, as in hydrocarbon stabilization, tray design is well understood, and many engineers are more comfortable with trays than with packing. In the past, bubble cap trays were the standard. However, they are not commonly used in this service anymore. Sieve trays are inexpensive but offer a very narrow operating range when compared with valve trays. Although valve trays offer wider operating range than sieve trays, they have moving parts and so may require more maintenance. High capacity/high efficiency trays can be more expensive than standard valve trays. However, high capacity/high efficiency trays require smaller diameter stabilization towers, so they can offer significant savings in the overall cost of the distillation tower. Random packing has traditionally been used in small diameter (<20 inches) towers. This is because it is easier and less expensive to pack these small diameter towers. However, random packed beds are prone to channeling and have poor turndown characteristics when compared with trays. For these reasons, trays were preferred for tower diameters greater than 20 inches.
Stripping Service
For stripping service, as in a glycol or amine contactor, bubble cap trays are the most common. In recent years, there has been a growing movement toward crimped sheet structured packing. Improved vapor and liquid distributor design in conjunction with structured packing can lead to smaller-diameter and shorter stripping towers than can be obtained with trays.
Figure 5-30. Various types of random packing.
Figure 5-31. Structured packing can offer better mass transfer than trays.
5.7 Glycol Dehydration System Sizing
Sizing Involve:
Glycol gas contactor diameter
Number of absorber trays (which establishes the tower’s overall height), tray spacing, and downcomers sizing
Glycol circulation rate
Lean glycol concentration
Reconcentrator heat duty
The number of absorber trays, glycol circulation rate, and lean glycol concentration are all inter-related.
5.7.1 Contactor Diameter
Method #1
The minimum diameter for trayed towers and conventional packing can be determined from the following equation:
d2 = 5040 [(T0ZQg)/P] [(ρg / ρL – ρg)(CD/dm)]0.5 Eq. 5-5
Where:
d = Contactor inside diameter, inches
dm = Drop size, microns =120 to 150 micron range
To = Contactor operating temperature, 0R
Qg = Design gas flow rate, MMscfd
P = Contactor operating pressure, psia
CD = Drag coefficient
ρg = Gas density, lb/ft.3 = 2.7 (SP/TZ) or = ρg= 0.093 ((MW)P)/TZ lb/ft3 (Eq. 1-19)
ρL = Glycol density, lb/ft.3 = 70 lb/ft3
Z = Compressibility factor
S = Gas specific gravity (air = 1)
Structured packing can handle higher gas flow rates for the same diameter contactor.
Method #2
Contactor diameter is set by the gas velocity. Sizing is calculated using recommended values for K-factors and C-factors are shown in table. 5-3, and equations 5-6
Table. 5-3. Recommended sizing parameters for TEG contactors.
G = C [ρV (ρL - ρV)]0.5 Eq. 5-6
Where
G = mass velocity, lb /(ft2 • hr)
C = constant from table 5-3,
ρV = Vapor density, lb/ft3
ρL = Liquid density, lb/ft3
where
2 theoretical stages ≅ 8 bubble cap trays @ 24 inch tray Spacing Eq. 5-7
2 theoretical stages ≅10 ft of structured packing Eq. 5-8
Note: Structured packing vendors frequently quote an Fs value for sizing glycol contactors, where Fs is defined in Eq 5-9.
Fs = v (ρv)0.5 Eq. 5-9
Where
v = vapor velocity, ft/sec
ρv= density, lb/ft3
Values of Fs = 2.5 to 3.0 will generally provide a good estimate of contactor diameter for structured packing.
Method #3
Figures 5-32 through 5-35 are correlations prepared by vessel manufacturers that allow graphical solutions of glycol gas contactor diameters.
Fig. 5-32 Determination of contactor diameter—Sivills.
Fig. 5-33 Determination of contactor diameter—Smith Industries.
Fig. 5-34 Determination of contactor diameter—NATCO.
Fig. 5-35 Determination of contactor diameter—BS&B.
5.7.2 Number of Trays and Tray Spacing
6 to 8 trays are used to meet normal dew-point depressions.
12 trays are typically required for high dew-point depressions.
Spacing ranges from 20 to 30 inches
24 inches is preferred, while 30-inch spacing is recommended if foaming is anticipated
5.7.3 Downcomers
Sized for a maximum velocity of 0.25 ft./sec.
5.7.4 Glycol Circulation Rate
For a given dew-point depression, the circulation rate is dependent upon:
Lean glycol concentration
Number of trays
When the lean glycol concentration and number of trays are held constant, the required glycol circulation rate can be determined from the following equation:
L = (ΔW/Wi) Wi Qg/24 Eq. 5-10
Where:
L = Glycol circulation rates, gal/hr
ΔW/ Wi = Circulation ratio, gal TEG/lb H2O (see Figures 5-36, through 5-41)
Wi = Water content of inlet gas, lb H2O/MMscf
W0 = Desired outlet water content, lbH2O/MMscf
ΔW = Wg / W0
Qg = Gas flow rate, MMscfd
5.7.5 Lean Glycol Concentration
Equilibrium water dew points for various concentrations of TEG are shown in Figure 5-10.
Glycol purity (lean glycol concentration) is a function of the temperature of the reconcentrator (Fig. 5-9).
Glycol purity can be increased by:
a- Adding stripping gas
b- Reducing the pressure in the reconcentrator
c- Reducing the glycol circulation rate
5.7.6 Glycol-Glycol Preheater
Cool wet glycol from the contactor enters the preheater (heat exchanger) at 100 0F and the warm glycol leaves at 1750 to 200 0F en route to the gas/glycol/condensate separator.
Hot dry glycol from the glycol/glycol heat exchanger enters the preheater at 250 0F and the warm dry glycol leaves at 150 0F to the glycol pumps en route to the contactor.
Temperature limitations to the glycol pump:
Glycol powered pumps (Kimray) limited to 200 0F.
Electric plunger pumps limited to 250 0F.
Overall heat transfer coefficient (U = 10 to 12)
5.7.7 Glycol-Gas Cooler
TEG to gas contactor is limited to 10 0F to 15 0F above the inlet gas temperature. If hotter, some TEG will vaporize with gas. If colder, gas condensation of the hydrocarbons may cause foam and glycol loss.
Overall heat transfer coefficient (U = 45).
5.7.8 Glycol-Glycol Heat Exchanger
Hot dry glycol from the reconcentrator enters the heat exchanger at 390 0F and leaves at 250 0F en route to the glycol/glycol preheater. Warm wet glycol from the charcoal filter enters the heat exchanger at 200 0F and the hot wet glycol leaves at 350 0F en route to the still column.
5.7.9 Gas-Glycol-Condensate Separator
Separator should be sized using procedures for sizing gas-liquid separation.
Liquid retention times between 20 and 30 minutes are recommended, depending on API gravity of the condensate. Operating pressure of 35 to 50 psig is recommended.
5.7.10 Reconcentrator
The reconcentrator should be designed to operate 350 0 to 400 0F with TEG, and 305 0F with DEG. Design temperature should be sufficiently below the decomposition point so that hot spots on the fire tube and poor mixing in the reconcentrator will not cause decomposition of the glycol. With everything else operating normally, the reconcentrator temperature is raised to lower the water content of the treated gas, and vice versa. Specific reconcentrator operating temperature is determined by trial and error.
Temperatures up to 400 0F are common
400 0F yields 99.5% TEG purity
375 0F yields 98.3% TEG purity
5.7.11 Heat Duty
Estimated from the following equations
qt = LQL Eq. 5-11
Where:
qt = Total heat duty on reconcentrator, Btu/hr
L = Glycol circulation rate, gal/hr
QL = Reconcentrator heat load, Btu/gal TEG (Table 5-4)
Heat duty estimated from Equation (5-11 above) is normally increased by 10 to 25% to account for start-up, fouling, and increased circulation.
Table. 5-4. Reconcentrator Heat Load
Alternatively, Heat duty can be determined as follows:
Total heat duty =
Sensible Heat, Qs + water vaporization heat duty, QV+ Condenser Duty, QC, Eq. 5-12
Qs = m Cp Δt Eq. 5-13
Where
Qs = Sensible Heat, Btu / gal
m = density lb/gal (table 5-1)
CP = Heat capacity, (Btu/lb.°F)
CP (95.1% TEG) = 0.56 at 110 0F (from physical property of TEG), 0.63 Btu/hr 0F at 200 0F, and 0.70 at 300 0F.- (use 0.67 as an average for the range of 200-300 0F).
Δt = Temperature difference (0F)
Vaporization of Absorbed H2O:
Qv = (ΔHvap) (ΔW) Eq. 5-14
Where
Qv = vaporization of water heat duty, Btu/gal.
ΔHvap = latent heat of vaporization, Btu/lb
ΔW =change of water content
Condenser Duty @ 25% Reflux Ratio:
Qc = % Reflux Ratio X Qv /100 Eq. 5-15
5.7.12 Fire Tube Sizing
The actual surface area of the firetube required for direct-fired heaters can be calculated from the following equation:
A = qt / 6000 Eq. 5-16
Where:
A = Total firetube surface area, ft.2
qt = Total heat duty on reconcentrator, Btu/hr
By determining the diameter and overall length of the U-tube fire tube, one can estimate the overall size of the reconcentrator. A heat flux of 6000 to 8000 Btu/hr-ft.2 is often used, but the 6000 value is suggested to ensure against glycol decomposition.
Fig. 5-36. Water Removal vs. TEG Circulation Rate at Various TEG Concentrations (N = 1.0)
Fig. 5-37. Water Removal vs. TEG Circulation Rate at Various TEG Concentrations (N = 1.5)
Fig. 5-38 . Water Removal vs. TEG Circulation Rate at Various TEG Concentrations (N = 2.0)
Fig. 5-39 . Water Removal vs. TEG Circulation Rate at Various TEG Concentrations (N = 2.5)
Fig. 5-40. Water Removal vs. TEG Circulation Rate at Various TEG Concentrations (N = 3.0)
Fig. 5-41. Water Removal vs. TEG Circulation Rate at Various TEG Concentrations (N = 4.0)
5.7.13 Reflux Condenser
Wet glycol inlet from the gas contactor enters at 115 0F and leaves at 125 0F.
It controls TEG losses. Reflux rate should be 50% of the water removal rate. The condensing coil provides lowest TEG loss and most economical reconcentrator operation.
5.7.14 Stripping Still Column
Temperature considerations:
Temperature is critical to the operation of the still column. Heat is provided by the reconcentration. Reconcentrator temperatures in the range of 350 0F to 400 0F insures adequate heat transfer to the ceramic packing in the still column.
Still columns whose wet glycol inlet enters above the packed section (Figure 5-42):
Operate best with a vapor outlet temperature between 2250 and 250 0F
Purpose of glycol falling over ceramic packing is the efficient use of available heat
Backpressure should be kept to a minimum (1 psig is maximum)
Still columns whose wet glycol inlet enters below the packed section (Figure 5-43):
Allow pall ring type packing to be solely involved in the reflux process
Operate best with a vapor outlet temperature between 185 0F and 195 0F
This temperature allows a greater volume of condensation by the reflux coil while still permitting the majority of the steam to escape
Fundamentals of Oil and Gas Processing
Basics of Gas Field Processing
Basics of Corrosion in Oil and Gas Industry
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Basics of Gas Field Processing
Basics of Corrosion in Oil and Gas Industry
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https://oilprocessing.net/oil/index.php
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Re: Basics of Gas Field Processing Book "Full text"
Gas Dehydration - Chapter 5 - Part 2
Still Column Diameter Size
Diameter size is based on the required diameter at the base of the still, calculated by vapor and liquid loading conditions at that point. Vapor load consists of the water vapor (steam) and stripping gas flowing up through the still. Liquid load consists of the rich glycol stream and reflux flowing downward through the still column. The diameter required for the still is based on the glycol circulation rate (Figure 5-44).
Still Column Packing
1 to 3 theoretical trays (4 to 12 feet) is sufficient for most TEG stripping still requirements.
304 SS packing is normally used.
Amount of Stripping Gas
The amount of stripping gas required to reconcentrate the glycol to a high purity will range from 2 to 10 ft.3/gal TEG circulated (Figure 5-45).
5.7.15 Filters
Microfiber : Sized to remove 5 micron solids.
Activated Charcoal (Carbon) :
Used to remove chemical impurities. Sized for full flow with 10 gpm streams.
Sized for 10 to 25% side streams on large units.
Fig. 5-42 Still column with wet glycol entering above the ceramic saddle packing.
Fig. 5-43 Still column with wet glycol entering below the stainless steel pall rings.
Fig. 5-44 Determination of stripping still column diameter.
Fig. 5-45 Amount of stripping gas required to reconcentrate glycol to high purity.
5.8 Calculation Examples for Glycol Dehydration
Example 5-1
30 MMscf/d of a 0.65 sp gr natural gas enters a TEG contactor at 600 psia and 100°F. Outlet water content specification is 7 lb H2O/MMscf and the TEG circulation rate is 3 gal TEG/lb H2O. Estimate the contactor diameter and number of bubble cap trays or height of structured packing required to meet this requirement. Assume z = 0.92.
Gas mol. Wt. = 0.65 X 28.96 = 18.8
Gas density, ρg =P (MW)/ RTZ
= (600) (18./ (10.73) (560) (0.92) = 2.0 lb/ft3
Liquid density “from table 4-4”, ρL = 1.119* (62.4) = 69.9 lb/ft3
From fig. 4-8. Water inlet (Win)= 90 lb/MMscf
Solutions Steps:
1. From Fig. 4-8 (McKetta and Wehe chart) is equivalent to a water content of 7 lb H2O/MMscf @ 600 psia)
2- Estimate required TEG concentration from Fig. 5-10, H2O Dewpoint = 24°F,
@ T = 100°F, lean TEG concentration ≈ 98.8 wt%
3. Estimate number of theoretical stages.
Calculate water removal efficiency
(Win −Wout) / Win = (90 −7) /90 = 0.922
From Fig. 5-37 (N = 1.5) at 3 gallon TEG/lb H2O and 99.0 wt% TEG
(Win – Wout)/Win = 0.885
From Fig. 5-38 (N = 2.0) at 3 gallon TEG/lb H2O and
99.0 wt% TEG
(Win – Wout)/Win = 0.925 use N = 2.0
Using equations 5-6, 5-7, and 5-8.
G = C [ρV (ρL - ρV)]0.5 Eq. 5-6
2 theoretical stages ≅ 8 bubble cap trays @ 24 inch tray Spacing Eq. 5-7
2 theoretical stages ≅10 ft of structured packing Eq. 5-8
“C” from table 5-3
4. Size the contactor Bubble caps, 24 inch tray spacing:
G = 576 [2.0 (69.9 −2.0]0.5 = 6712 lb /ft2.hr
Mass flow, m =
M = 62000 lb/hr
Area of flow (A) = m “mass flow” /G “mass velocity” = 62000 / 6712 = 9.2 ft2
Since, area ft2 = π D2 /4
D2 = 9.2 X 4 / 3.14 = 11.7
D = 3.4 ft.
Alternatively
using equation 5-5 ( drag coefficient CD =0.85 and dm = 150 micron)
d2 = 5040 [(T0ZQg)/P] [(ρg / ρL – ρg)(CD/dm)]0.5 Eq. 5-5
d2 = 5040 [(560 x 0.92 x 30 x 106)/600] [(2 / 67.9)(0.85/150)]0.5
minimum diameter D = 3.5 ft.
5. Size For Structured packing: ( “C” from table 5-3 “ use 1260 as an average”)
G = 1260 [2.0 (69.9 −2.0]0.5 = 14683 lb /(ft2.hr)
Mass flow, m = (30 X106) (0.65 X 28.96) / (379.5) (24) = 62000 lb/hr
Area A = m/G = 62000 / 14683 = 4.22 ft2
Area ft2 = π D2 /4
D2 = 4.22 X 4 / 3.14 = 5.38
D = 2.3 ft
Example 5-2
Determine reboiler duty for conditions in the previous example. Assume the rich TEG temperature entering the regenerator is 300°F and the reboiler temperature is 400°F.
Glycol Reboiler Duty: Basis 1 gal. TEG. ( assume 25% reflux ratio)
Solution:
Using equation 5-12, and 5-13
Qs = m Cp Δt Eq. 5-13
- (use 0.67 as an average for the range of 300-400 0F)
Qs = (9.3 lb/gal.) X (0.67 Btu/lb.°F) X (400°F − 300°F) = 623 Btu /gal.
Using eq. 5-14
Qv = (ΔHvap) (ΔW) Eq. 5-14
Qv = (970 Btu/lb H2O) X (1 lb H2O/ 3 gal. TEG) = 323 Btu / gal.
Using eq. 5-15
Qc = % Reflux Ratio X Qv /100 Eq. 5-15
Qc = (25) (Qv)/100 = 81 Btu /gal.
From eq. 5-12
Total Duty Including 10% Heat Loss:
Qr (total regeneration heat duty) = (623 + 323 + 81) (1.1) = 1130 Btu /gal.
Total Duty Based on 30 MMscfd of Gas:
Q = (1130 Btu/gal) (3 gal./lb) (30 MMscfd/24) ((90-7) lb/MMscf) = 350,000 Btu/hr
Example 5-3
Given: Gas Qg = 98 MMscfd at 0.67 SG saturated with water at 1000 psig and 100 0F
Dehydrate to = 7 lb/MMscf
Use triethylene glycol, No stripping gas is available
98.5% TEG concentration
CD (contactor) = 0.852
Z = 0.865
Determine:
1. Calculate contactor diameter
2. Determine glycol circulation rate.
3. Size the still column
Solution:
1. Calculate contactor diameter
d2 = 5040 [(T0ZQg)/P] [(ρg / ρL – ρg)(CD/dm)]0.5 Eq. 5-5
dm = 125 microns, T = 570 0R
P = 1015 psia, Qg = 98 MMscfd
ρL = 70 lb/ft.3
ρg = (0.67 X 1015)/(560 X 0.865) = 3.79 lb/ft3
d2 = 5040 [(560 X 0.865 X 98)/1015] [(3.79/70-3.79)(0.852/125)]0.5
d = 68.2 in.
Use 72.00 ID contactor (standard off-the-shelf)
2. Determine glycol circulation rate and reboiler duty
Wi = 63 lb/MMscf (from McKetta-Wehe) “saturated water content”
W0 = 7 lb/MMscf (spec)
ΔW = Wi - W0 = 63 - 7 = 56 lb/MMscf
ΔW/Wi = 56/63 = 0.889
Using n = 2 (i.e., 8 actual trays) and glycol purity of 98.5% read from Figure 5-38 the glycol circulation rate of about 2.7 gal TEG/lb H20. Use 3.0 gal/lb for design.
Using equation
L = (ΔW/Wi) Wi Qg/24 5-10
L = (3.0 gal/ Lb) (56 lb/MMscf)(98 MMscf/day)(day/24 hr)(hr/60min)
= 11.4 gpm TEG
= 862 Btu/gal (From table 5-4)
= (862 Btu/gal) (11.4 gal/min) (60 min/hr)
= 590,000 btu/hr
To allow for start-up heat loads, increase heat duty by 10% and then select a standard off-the-shelf fire tube.
Thus, select a 750 MMBtu/hr.
3. Design of still column:
Use 12-foot still column (standard packed arrangement)
dm = 125 micron
T = 300 0F = 760 0R
P = 1 psig = 15.7 psia
Qg = (10 scf/gal) (11 gal/min) (60 min/hr)(24 hr/day)
= 0.16 MMscfd
Z = 1.0
Since ρg = Gas density, lb/ft.3 = 2.7 (SP/TZ) or = ρg= 0.093 ((MW)P)/TZ lb/ft3 (Eq. 1-19)
ρg = 2.7 (0.62 X 15.7)/ (760 X 1.0 )
= 0.035 lb/ft.3
ρL = 62.4 lb/ft.3
CD = 14.2 (given)
d2 = 5040 [(760 X 1 X 0.16)/15.7] [(0.035/62.4-0.035)(14.2/125)]0.5
d = 17.7 in.
Use 18 inch OD x 12 feet long still.
5.9 Glycol Unit Operation
5.9.1 Start up
Prior to the initial start up of a new plant, the vessels and lines should be thoroughly washed out with water to remove debris and corrosion products that accumulated during construction. After the system has been cleaned, start up is accomplished in three phases:
Establish glycol circulation throughout the plant.
Apply heat to the reboiler and bring it up to operating temperature.
Open the wet gas stream to the contactor and begin dehydrating the gas.
In order to circulate glycol throughout the system, it will be necessary to pressurize the vessels in the system. Pressuring can be done with wet gas or dry gas. The contactor pressure should be raised to at least 150 psig (10 bars) and the flash tank pressure should be raised to at least 45 psig (3 bars). When the vessels have been pressured, start up procedure is:
Fill the reboiler and surge tank with fresh glycol solution Also add to the flash tank.
Pressure up the contactor column by very slowly opening the gas inlet valve.
Prime and start glycol pump.
When liquid appears at the bottom of the contactor put the bottom level controller in service so the glycol will flow to the flash tank.
Put the flash tank level controller in service when liquid appears in the bottom, so that liquid will flow to the stripper.
Keep surge tank level half full by adding glycol when needed.
When desired circulation rate is established, light the reboiler or put the heat source in service and slowly bring reboiler temperature up to 250°F (121°C). Leave temperature at 250°F (121°C) until all water has been boiled out of glycol.
5.9.2 Routing Operation
Routing operating checks include the following:
Check levels in each vessel and reset level controller as necessary.
Check the pressure drop across the filter and replace the elements as required.
Check the temperature of the lean glycol out of the glycol exchanger to see that the proper transfer rate is occurring in the exchanger.
Check the flow of glycol to the contactor and of stripping has to the reboiler.
Check the pressure of the flash tank to see if it is at proper level.
If water or air is used to cool the glycol prior to its entry into the contactor, check the glycol temperature in order to ensure that it is about 5°C to 7°C (10°F to 15°F) above the inlet gas temperature. Adjust the flow of air or water through the cooler as required.
Check the water content of the outlet gas to see that it is below the design limit.
5.9.3 Shut Down
This procedure is used to shut down a glycol plant:
Block gas to contactor column,
Shut off heat (leave pump running),
When unit cools to safe temperature, less than 200°F, shut off pump.
Drain glycol, if necessary.
5.10 Glycol Maintenance and Care
5.10.1 Preventive Maintenance
Scheduled preventative maintenance reduces glycol losses and operating problems such as foaming, system plugging, corrosion, pump failures.
It also minimizes system down time and maximizes system operation efficiency.
Five Steps to a Successful Preventive Maintenance Program
Record-Keeping
Accurate records can be used to determine the system efficiency and to pinpoint operating problems. Records of prior and existing conditions including dew points, glycol usage, and repairs help establish the system profile. Once the system profile is defined it becomes easier to identify unusual system characteristics that may indicate potential problems.
Mechanical Maintenance
Daily physical inspections are necessary to insure that the system is running properly. Any trouble encountered should be dealt with immediately, thus preventing the problem from escalating.
Glycol Care
Regular chemical analysis (every one or two months) of the glycol provides detailed information on the internal operation of the unit. Many process-related problems can be diagnosed well in advance of mechanical failure. Chemical problems can be diagnosed and corrective action taken before they become costly and detrimental to unit performance.
Corrosion Control
Corrosion is a frequent problem in glycol dehydration systems. If unchecked the damage can be extensive. All units should have provisions for corrosion control.
Communication
Lines of communication between field and technical assistance or office personnel are critical to the smooth operation of any system. Office personnel (production supervisors, facility engineers, purchasing agents) must be kept informed of daily operations and any problems that may arise.
Field personnel must be made aware of technical information that may improve their operations. Training for field operators allows the operator to better maintain the equipment.
5.10.1.1 Record-Keeping
Records necessary to establish a system profile include:
Design information including vessel specifications, equipment drawings, and P&IDs
Filter element or media replacement—type and frequency
Glycol usage—gallons/month
Chemical additives—type and amount
Gas production and flow rate charts—peak, average, and low periods
Outlet gas dew point/water content (lbs/Mscfd)
Mechanical inspections—type, magnitude, frequency, results
Records necessary to establish a system profile include:
Glycol analysis—format, frequency, recommendations, results
Corrosion coupon results—mills per year (MPY), frequency
Materials and labor relating to system repairs— operating costs
With the aforementioned information, a good system profile can be drawn of a specific system.
Updating these records will show any gradual changes in a unit’s system profile and may alert you to a potential or developing problem.
5.10.1.2 Mechanical Maintenance
The following things should be done so as to keep the unit operating properly and to prevent operational problems:
1. Insure that instruments and controls are in good working condition (thermometers and pressure gauges, etc.). Use a test thermometer on the reconcentrator to insure proper reconcentrator heat.
2. Insure glycol filter elements are changed according to average expected need basis:
Microfiber filters should be changed monthly.
Carbon filters should be changed monthly (small cartridge filters) to every six months (large bulk units). Glycol analysis helps determine the frequency.
An upset or sudden change in the operating conditions may foul the filters faster than the preventive maintenance program anticipates. Make sure filter differential pressure is below 15 psi.
3. Look for glycol leaks on and around the glycol skid. Most leaks can be stopped by tightening a union, valve stem packing, or pump rod packing. After the leak has been repaired, clean the affected area so it is easier to notice new leaks.
4. Check the glycol level, at least twice a day, and add glycol as necessary. Maintain a written report of glycol added. This allows operations to detect excessive losses of glycol and take corrective action faster.
5. Insure unit performance by taking a dew point measurement daily.
6. Clean the glycol strainers monthly to prevent accumulation of trash, which can cause the glycol pump to fail.
7. Check the glycol circulation rate daily. Any time the gas flow rate changes or when a drastic change in gas pressure or temperature is experienced, the glycol flow rate should be recalculated and the pumps set accordingly. On multiple pump installations, switch the pumps weekly, thus insuring pump operation when necessary.
8. On direct fired fire tubes, on weekly basis; sight down the flame direction, to insure it is not touching the fire tube for fire tube blisters or hot spots.
9. Cycle the main burner manually to be sure the fuel gas valve works and the pilot light stays lit. Check the fuel gas scrubber pot for fluid build-up that may hinder burner operation.
5.10.1.3 Glycol Care
Operating and corrosion problems occur when the circulating glycol gets dirty. Some contaminated glycol problems can be noticed easily and corrective action taken.
A small glycol sample should be taken daily from the surge tank or dry glycol suction header to the pump. Check closely for fine black particles settling out of the sample, which may be corrosion by-products and indicate an internal corrosion problem.
Uncontaminated glycol is colored like pure kitchen oil. Glycol gets darker and darker due to hydrocarbon contaminates, and/or corrosion products.
Smell the sample, If the smell is sweet and aromatic (similar to rotten bananas) it may be thermally decomposed. If the sample is viscous and black it is probably contaminated with hydrocarbon or well-treating chemicals. If the hydrocarbon contamination is great enough, the sample will separate into two liquids or interphase. Every one or two months send a sample of both the rich and lean glycol to a laboratory for complete analysis. This type of analysis will provide a detailed description of unit performance and glycol condition.
5.10.1.4 Corrosion Control
Corrosion is a major cause of premature equipment failure. Corrosion can occur over the entire system, inside and out. The two most common areas of severe corrosion are:
Still column reflux coil, and vent/fill connection on the surge tank.
This is due to a high concentration of water vapor in the top of the still and the ready availability of oxygen in the air at the vent/fill cap. Three types of corrosion are usually found in glycol systems either individually or in combination with one another are:
Oxidation
Is the corrosion process involved the presence of oxygen molecules. Some metal loss is incurred and the resultant corrosion product is a scale-like residue called oxide, or rust.
Oxidization is characterized by rough, irregular, shallow pitting of the metal scaled over the rust.
Sour Corrosion
Hydrogen sulfide (H2S) are often found in produced natural gas. Glycols are very reactive with sulfur compounds, such as H2S. The resulting materials tend to polymerize (form larger molecules) that form a “gunk” that is very corrosive.
Corrosion in the presence of acid gases is characterized by deep, jagged pitting.
Sweet Corrosion
Water is found in glycol as vapor, free condensed water, or entrained water in glycol. Carbon dioxide (CO2) when dissolved in water forms carbonic acid. Since most produced natural gases contain some CO2, the presence of carbonic acid in glycol systems is very common. The corrosion resulting from carbonic acid is known as sweet corrosion. Sweet corrosion is characterized by deep, round, smooth pitting. Sometimes the pitting will cover a broad area, disguising the depth of the pit.
5.10.1.4.1 Corrosion Monitoring and Control Programs
Monitoring and control programs should include system monitoring through:
Corrosion coupons
Glycol analysis (pH and iron)
Three steps in combating corrosion in glycol systems are:
Use an effective corrosion inhibitor in both the liquid and vapor phases.
Use corrosion resistant alloys (CRA) in construction.
Keep the unit clean to prevent acid formation due to contamination.
Cathodic protection has been attempted but met with little success.
It is impractical to attempt to eliminate corrosion. The rate of corrosion can be slowed to a point that is almost negligible. The maximum acceptable corrosion rate is 6 mils per year (MPY).
Corrosion inhibitors work in several ways.
The two most applicable to glycol units are:
pH buffers which include:
Alkanolamines
MEA (Monoethanolamine)
TEA
The pH buffers fight corrosion by stabilizing the pH near neutral, thereby reducing corrosive environment. Alkanolamines are regenerable as is the glycol and thus can be retained in the system for lengthy time periods. However, they are thermally degraded at normal reconcentrator operating temperatures and if used frequently may leave harmful residues within the system.
Plating Inhibitors
Tallow diamine, unlike the inorganic amines, is an organic amine. It is grouped with the plating inhibitors even though it does not actually plate out on the vessel walls. It flashes out of the glycol at high temperatures. As it vaporizes it contacts the vapor spaces of the reconcentrator and forms a tenacious film over the exposed metal. These inhibitors are strictly liquid phase protection. They will plate out on the vessel walls forming a protective barrier between a corrosive environment and the metal. This film will eventually wear away and must be replenished occasionally to continue protection.
Since the plating inhibitors are all alkalies, some degree of pH buffering will be effected.
The pH buffering will not be as great as through the use of amines.
True plating inhibitors include:
Borax
NaCap (Sodium mercaptobenzothiazole)
Dipotassium phosphate
5.10.1.5 Communication
Communication is the easiest portion of an effective maintenance program and yet it is the most overlooked. Communication can be between management and labor, engineer and foreman, operators on opposite shifts, and office and field personnel. Lack of communication is the single most contributing factor to glycol system failure. Failure to communicate can cause confusion and evolve into major problems.
5.11 Glycol Operation Considerations
Operating and corrosion problems usually occur when the circulating glycol gets dirty.
To achieve a long, trouble-free life from the glycol, it is necessary to recognize these problems and know how to prevent them.
Some of the major areas are:
Oxidation
Thermal decomposition
pH control
Salt contamination
Hydrocarbons
Sludge
Foaming
5.11.1 Oxidation
Sometimes glycol will oxidize in the presence of oxygen and form corrosive acids. Oxygen enters the system with the incoming gas through:
Unblanketed storage tanks and sumps
Pump packing glands
To prevent oxidation:
Bulk storage tanks should be gas blanketed.
Use oxidation inhibitors.
Normally, a 50/50 blend of MEA and 33.5% hydrazine is inserted into the glycol between the absorber and the reconcentrator. A metering pump should preferably be used to give continuous, uniform injection.
5.11.2 Thermal Decomposition
Excessive heat, a result of the following conditions, will decompose glycol and form corrosive products:
High reconcentrator temperature above the glycol decomposition level
High heat-flux rate, sometimes used by a design engineer to keep the heater cost low
Localized overheating, caused by deposits of salts or tarry products on the reconcentrator fire tubes or by poor flame direction on the fire tubes
5.11.3 pH Control
pH is a measure of the acidity or alkalinity of a fluid, based on a scale of 0 to 14. pH values from 0–7 indicate the fluid is acidic. pH values from 7–14 indicate the fluid is alkaline.
To obtain a true reading, glycol samples should be diluted 50-50 with distilled water before pH tests are run. The pH meter should be calibrated occasionally to keep it accurate.
New glycol has a neutral pH of approximately 7. With usage the pH decreases and the glycol becomes acidic and corrosive unless pH neutralizers or buffers are used. Equipment corrosion rate increases rapidly with a decrease in the glycol pH. Acid created by glycol oxidization, thermal
decomposition products, or acid gases picked up from the gas stream are the most troublesome of corrosive contaminants. In addition, a low pH accelerates the decomposition of glycol.
Ideally, the glycol pH should be held at a level between 7.0 and 7.5. A value above 8.5 tends to make glycol foam and emulsify. A value below 6.0 corresponds to system contamination, corrosion, and/or oxidation.
Borax, ethanolamines (usually triethanolamine) or other alkaline neutralizers are used to control the pH. These neutralizers should be added slowly and continuously for the best results.
An overdose will usually precipitate a suspension of black sludge in the glycol. The sludge could settle and plug the glycol flow in any part of the circulating system.
Frequent filter element changes should be made while pH neutralizers are added.
The amount of neutralizer to be added and the frequency will vary from location to location. Normally, 1/4 lb of triethanolamine (TEA) per 100 gallons of glycol is sufficient to raise the pH level to a safe range.
When the glycol pH is extremely low, the required amount of neutralizer can be determined by titration. For best results, the lean rather than the rich glycol should be treated. It takes time for the neutralizer to mix thoroughly with all the glycol in the system. Several days are required before the pH is raised to a safe level. Each time that neutralizer is added, the pH of the glycol should be measured several times.
5.11.4 Salt Contamination and Deposits
In areas where large quantities of brine are produced, some salt contamination will occur. Salt deposits accelerate equipment corrosion. It also reduces heat transfer in the fire tubes. It alters specific gravity readings when a hydrometer is used to measure glycol water concentration. It cannot be removed with normal regeneration. A scrubber installed upstream of the glycol plant should be used to prevent salt carry-over from produced free water. The removal of salt from the glycol solution is then necessary.
5.11.5 Hydrocarbons
TEG will typically absorb about 1 scf of sweet natural gas per gallon of glycol at 1000 psia and 100°F. Solubilities will be considerably higher if the gas contains significant amounts of CO2 and H2S. Heavier paraffin hydrocarbons are essentially insoluble in TEG. Aromatic hydrocarbons, however, are very soluble in TEG, and significant amounts of aromatic hydrocarbons may be absorbed in the TEG at contactor conditions. This may present an environmental or safety hazard when they are discharged from the top of the regenerator.
Vapor-liquid equilibrium constants (K-values) for benzene, toluene, ethylbenzene, and o-xylene in TEG solutions, indicates that at typical contactor conditions approximately 10-30% of the aromatics in the gas stream may be absorbed in the TEG solution.
Aromatic absorption increases with increasing pressure and decreasing temperature. Aromatic absorption is directly related to TEG circulation rate. Higher circulation rates result in increased absorption. Aromatic absorption is essentially independent of the number of contacts in the absorber so one method of minimizing aromatic absorption is to use taller contactors and minimize TEG circulation rates.
Most of the aromatic components will be stripped from the TEG solution in the regenerator.
Flash tank sizing should be sufficient to degas the glycol solution and skim entrained liquid hydrocarbons, if necessary. A minimum retention time of 3-5 minutes is required for degassing. If liquid hydrocarbons are to be removed as well, retention times of 20-30 minutes may be required for adequate separation. Flash tank pressures are typically less than 75 psia.
Regenerator sizing requires establishing the reboiler duty and, when high TEG concentrations are required, providing sufficient stripping gas.
Liquid hydrocarbons, a result of carry-over with the incoming gas or condensation in the contactor, affects glycol by foaming, degradation, and losses. It must be removed with glycol/gas/condensate separator, or hydrocarbon liquid skimmer, or activated carbon beds
5.11.6 Sludge
Solid particles and tarry hydrocarbons (sludge) are suspended in the circulating glycol, and with time will settle out (It looks like asphalt or paraffin sludge). This results in the formation of black, sticky, abrasive gum that can cause trouble in pumps, valves and other equipment, usually when the glycol pH is low.
5.11.7 Foaming and defoamers
Excessive turbulence and high liquid-to-vapor contacting velocities usually cause the glycol to foam (this condition can be caused by mechanical or chemical problems).
The best way to prevent foaming is proper care of the glycol, such as:
Effective gas cleaning ahead of the glycol system
Good filtration of the circulating solution
Defoamers serve only as a temporary control until the conditions generating foam can be identified and removed. Success depends on when and how it is added. Most are inactivated within a few hours under high temperature and pressure, and thus their effectiveness is dissipated by the heat of the glycol solution. Thus, defoamers should be added continuously, a drop at a time, for best results.
The chemical feed pumps should meter the defoamer accurately, improve dispersion into the glycol, and may be activated automatically by differential pressure across the contactor.
5.12 Analysis and Control of Glycol
Analysis of glycol is essential to good plant operation. It helps pinpoint high glycol losses, foaming, corrosion, and other operating problems. Analyses enable operations personnel to evaluate plant performance and make operating changes to obtain maximum drying efficiency.
5.12.1 Visual Inspection
A glycol sample should first be visually inspected to identity some of the contaminants.
A finely divided black precipitate may indicate the presence of iron corrosion products.
A black, viscous solution may contain heavy, tarry hydrocarbons.
The characteristic odor of decomposed glycol (a sweet aromatic odor) usually indicates thermal degradation. A two-phase liquid sample usually indicates the glycol is heavily contaminated with hydrocarbons. The visual conclusion should next be supported by a chemical analysis.
5.12.2 Chemical Analysis
A complete glycol analysis of lean and rich samples, when properly interpreted, can provide a detailed picture of the workings of the dehydration unit and its process.
Glycol analysis should include tests to determine the following (table 5-5):
Test Lean Glycol Rich Glycol Allowable Range Ideal
pH (50/50) 6 to 8 7 to 7.5
Hydrocarbon
(% wt.) 0.1%
Water content
(% wt.) 2% lean - 6% rich
TSS (% wt.) 0.01%
Residue (% wt.) 4% 2%
Chlorides (mg/l) 1500 1000
Iron (mg/l) 50 35
Foam character:
Height (ml) 20 to 30 ml
Stability (sec) 15 to 5 sec
Specific gravity 1.118 to 1.126
Table 5-5.Glycol analysis.
5.12.3 Chemical Analysis Interpretation
5.12.3.1 pH
A pH below 6 generally corresponds with system contamination, corrosion, and/or oxidation.
Below 5.5 autoxidation occurs. Where Chemical decomposition of the glycol occurs within itself, and glycol has the tendency to continue to drop without outside influences.
The causes of low pH are:
Acid gases in the gas stream
Organic acids due to oxidation or thermal degradation
Excessive chlorides (salt) in the glycol
Well-treating chemicals entrained in the gas stream
Thermal decomposition of entrained liquid hydrocarbons in the gas stream and glycol
Oxidation of the glycol due to improper storage
While the causes of high pH are:
Contamination from well-treating chemicals entrained in the gas stream
Overdose of neutralizer added to a system for low pH
Foaming tendencies can result from high pH, due to stabilized glycol-hydrocarbon emulsions.
Sludge and residue build-up can result from both high and low pH.
Glycol pH should be checked periodically and kept on the alkaline side by neutralizing the acidic compounds with alkaline chemicals, such as monoethanolamine (MEA). A pH of about 7.3 is considered a safe level. Raising the pH above 8 to 8.5 is not desirable because of the tendency for an alkaline glycol solution to foam and emulsify more easily.
5.12.3.2 Sludge
Sludge may become abrasive and cause premature pump and valve failure. It may deposit in trays and downcomers, still column packing, and heat exchangers, which cause system plugging.
5.12.3.3 Hydrocarbons content
Enter the glycol stream as a result of inlet separator carry-over or as condensation due to temperature variations. Compressor lube oils and other extraneous organic chemicals such as pipeline corrosion inhibitors, are often stripped out of natural gas as it passes through the contact tower. Oils and organic residues can cause glycol/water emulsions and suspensions, which contribute to foaming, which will results in excessive high glycol carryover from the contactor, and the contaminants may cause plugging in the contactor, still column, and heat exchangers
The light hydrocarbons are usually separated from the glycol stream with an adequately sized glycol/hydrocarbon separator, while heavy hydrocarbons that are referred to as soluble hydrocarbons because they bond with the glycol are usually filtered out with activated carbon.
Light end hydrocarbons (insoluble) are allowable up to 1% by volume.
Soluble hydrocarbons are only acceptable to 0.1 % by weight, since they are primarily responsible for foaming, sludge and residue build up, low pH, loss of hygroscopicity, and glycol decomposition.
Hydrocarbons are often get into glycol in;
Condensation which is caused when glycol enters the contactor colder than the incoming gas. This problem can be eliminated by maintaining a temperature for the entering lean glycol of 10-15°F (5-7°C) warmer than the incoming gas.
Carryover of hydrocarbon contaminants from the inlet separator or gas
5.12.3.4 Water Content
Water content is defined as the quantity of water in the glycol. The difference between the lean sample and rich sample measures the degree of loading in the contactor. It indicates regeneration efficiency. Glycol purity should be at least 98% in the lean stream and at least 94% in the rich. These concentrations will produce the desired dew points in systems that are operating properly. For lower dew points the glycol purity must be increased (or water content decreased). High water content of the lean sample generally indicates low reconcentrator heat or one of the following reasons:
Excessive glycol circulation
Undersized equipment
Carry-over from the separator
Vapor communication from reconcentrator to surge
A leak in the glycol/glycol heat exchanger
Over-refluxing in the still column
Hot inlet gas temperature
High water content in the rich sample usually indicates a low glycol circulation rate or:
Carry-over from the separator
Poor reconcentration
Heat exchanger communication
Undersized equipment
Hot inlet gas temperature
Check values for hydrocarbon, chlorides, iron, and foaming to help pinpoint the problem.
5.12.3.5 Suspended Solids
Considered to be those solids and tarry hydrocarbons that remain suspended within the glycol solution down to 0.45 micron in size. They are result of poor inlet separation, corrosion, and thermal degradation of the glycol. Values greater than 0.01% by weight indicate poor sock/microfiber filtration. Most filters are sized to remove particles to a size of 5 microns.
Particles larger than this in excessive amounts may serve to stabilize foaming tendencies in glycol. When the glycol is allowed to maintain a large concentration of suspended solids, a silty residue is likely to form along vessel walls causing plugging of the contractor trays, heat exchangers, still column, and reconcentrator. Suspended solids are likely (common with low glycol pH).
Problems resulting from a high solids content include:
Increased pump wear from abrasion
Accelerated corrosion and erosion
Increased fouling of fire tube .
Increased glycol loss due to foaming
Increased plugging problems.
5.12.3.6 Residue
The value for residue is a function of system contamination. The glycol sample is distilled, removing all light end hydrocarbons, water, and glycol. Residue represents the remaining contamination, which is comprised of total solids (suspended and residual), salt, and heavy hydrocarbons.
Value for residue is best kept below 2% by weight, however some systems may operate reasonably well at values from 2% to 4%. Units with Glycol containing greater than 4% are prime candidates for failure and should be cleaned immediately.
5.12.3.7 Chlorides
Chloride values indicate the quantity of inorganic chlorides (salts) found in the glycol sample.
As the concentration of chlorides (as NaCI or CaCI) in glycol increases, its solubility decreases.
When heat is added to the glycol solution, the salt begins to form crystals which:
Fall out of the glycol solution
Accumulate on the heat source and can lead to premature heat tube failure
May be swept by the glycol into other areas of the system
Potential problems with excessive chlorides include system plugging, low pH, glycol pump damage, foaming, and loss of hygroscopicity due to rapid glycol decomposition.
Removal of chlorides in high concentrations requires vacuum distillation of the glycol.
Concentrations greater than 1000 ppm will stabilize foaming tendencies, may lead to excessive glycol loss, and may affect glycol pH.
Precipitation of salts from the glycol will begin at approximately 1200 to 1500 ppm, however, the crystals formed are extremely small and rarely troublesome. At concentrations above 2200 ppm, precipitation occurs readily and system failure is a possibility.
Filtration removes large salt crystals, but most of the damage associated with salt will have already occurred prior to the development of crystals large enough to filter.
5.12.3.8 Iron
Iron found in glycol samples can indicate possible corrosion, and/or produced water carry-over.
Iron in excess of 50 ppm generally indicates corrosion. Whether it be in the glycol unit, upstream in the production equipment or downhole in the well string is difficult to determine.
Comparing values for iron content in several points downstream the glycol unit helps to establish the location of suspected corrosion. Corrosion by-products will consist of soluble iron and fine, gritty particulate in systems where oxygen is available. In systems where no oxygen is present, corrosion by-products will include sulfides in addition to the iron.
5.12.3.9 Foaming
More glycol is lost through foaming than any other cause. Foaming is not easily detected without chemical analysis; gradual low-volume glycol loss often goes overlooked. It is usually a result of contamination. Primary contaminants that cause foaming are hydrocarbons (from separator carryover), suspended solids, chlorides, compressor lube oil, well treating chemicals, and iron.
Water content affect foaming tendencies by inducing emulsification of contaminants, particularly hydrocarbons. Carbon filtration is the most effective means of controlling foam. Silicone emulsion–type foam inhibitor is used, but they treat the symptom, not the cause and thus are temporary solutions. Addressing the source of the contamination causing the foam is the only long-term solution.
The foam test consists of bubbling dry air at a rate of 6 liters/min through a graduated cylinder container of 200 mm of the glycol sample until the foam stabilizes at its maximum height.
Volume for both the liquid and the foam is reported as a single value. The original 200 ml is then subtracted. The remaining value is recorded as height and represents the ease at which the solution will foam. Once the maximum foam height is recorded, the dry air is removed from the sample and the time it takes for the foam to break from its maximum volume to a clear surface on the glycol sample is recorded in seconds. This time represents the tendency of the foam and is known as stability.
There are no concrete values given for acceptable foam height and stability. Foam with very low height and moderate stability will result in little glycol loss as will a foam with moderate height and very low stability. Thus, the acceptable range for foam test results are:
Height/ml: 20 to 30 ml
Stability/sec: 15 to 5 sec
For example, a sample with a height of 25 ml and a stability of 10 sec is acceptable, while a sample with 30 ml height and 15 sec stability would have a high foaming tendency and could result in glycol losses.
5.12.3.10 Glycol Weight Percentage
This refers to the amount of glycol in the glycol solution. Lean glycol should contain about 98.5 to 99.9% a glycol. Rich glycol content varies from about 93-96% glycol.
5.12.3.11 Specific Gravity
Specific gravity is used to determine the purity of glycol. A specific gravity of 1.126 to 1.128 at 60 0F indicates a 99% TEG (technical grade).
A specific gravity of 1.124 to 1.126 indicates 97% (industrial grade).
With glycol extracted from an operating dehydration unit, the lean sample should have a specific gravity of 1.1189 to 1.121.
This variance allows for acceptable amounts of system contamination.
Low specific gravity would indicate one or more of the following:
TEG containing excessive amounts of EG and/or DEG (poor quality replacement glycol)
Excessive water in sample
Excessive hydrocarbons in sample
A high specific gravity indicates one or more of the following:
The system is contaminated with excessive amounts of solids or any additives with a greater density than glycol
Thermal degradation of the glycol
Oxidation or chemical degradation of the glycol
5.12.3.12 Glycol Composition
The composition of glycol indicates its quality.
Values are given to the component glycols (EG, DEG, TEG, TTEG) contained within the glycol sample solution. Industrial grade (97%) TEG or better is required for best glycol system results.
In addition to 97% TEG, the glycol solution may contain, in various concentrations, up to 1% EG and 3% DEG, but not to exceed a combined total of 3%.
Glycol degradation will often be reflected by changes in the glycol composition and reduction in pH. Thermal degradation is most common and is characterized by excessive values of EG, DEG, and occasionally the presence of TTEG. The thermal degradation is characterized by:
The glycol pH will be low.
The glycol sample will be dark and have an aromatic smell (ripe bananas).
Chemical degradation is brought about by oxidation and acidic contaminants and is characterized by:
Excessive values for EG and DEG but no TTEG will be present
Low pH
Glycol may not appear to be too dirty
Autoxidation is a form of continuing chemical degradation.
5.13 Troubleshooting
5.13.1 General Considerations
The most obvious indication of a unit malfunction is high water content (dew point) of the outlet stream. High water content is brought about by:
Insufficient glycol circulation
Insufficient reconcentration of the glycol
These problems can be caused by a variety of contributing factors such as mechanical causes or existing operating conditions for which the equipment was not designed
5.13.2 Main approach to troubleshooting:
Timeframe
Determine the approximate date/time at which the problem became apparent.
List Changes
Inventory any changes (things that happened differently than usual). Look for what is different.
Production changes
Operational changes
Maintenance and Repairs
Weather
Investigate by process of elimination reduce the list of changes to determine the factor or factors that manifest the problem.
5.13.3 High Dew Points
5.13.3.1 Insufficient Glycol Circulation
If there is insufficient glycol circulation, check heat exchangers and glycol piping for restrictions or plugging.
On an electric driven piston pump:
Check flow rate indicator (if present) to insure proper glycol circulation. If flow rate indicator is not present, verify circulation rate by closing the glycol discharge valve from the contactor and timing the fill rate in the gauge column.
Check high-pressure dry-glycol bypass valve. Close if necessary.
Check pump prime by shutting pump down, closing the discharge valve, opening the bypass valve and restarting the pump. Allow to run briefly under no load through the bypass line to remove any trapped gas in the pump.
On glycol-gas powered pumps:
Close dry discharge valve. If pump continues to run, open dry discharge bleed valve and allow running a few strokes. Once all gas is purged from put, close the bleed valve. If pump continues to run, discontinue use and send in for repair. If pump will not prime, but continues to run gas through the dry discharge bleed valve then:
Check pump suction strainer for plugging.
Check glycol level in surge tank.
5.13.3.2 Insufficient Reconcentration
Verify reconcentration temperature with test thermometer (3500 to 400 0F). Raise temperature if necessary.
Check glycol-to-glycol heat exchanger for leakage of wet glycol into the dry glycol stream.
Check stripping gas if applicable. Be sure stripping gas is in service at the proper rate.
Check for communication between the reconcentration vapor space and the surge tank vapor space.
5.13.3.3 Operating Conditions Different from Design
Check operation of upstream separators and scrubbers. Be sure not to overload system.
Increase absorber pressure. This may require installation of a back pressure valve.
Reduce gas temperature, if possible.
Increase circulation rate, if possible.
Increase reconcentrator temperature, if possible.
5.13.3.4 Low Flow Rate
Blank off a portion of the bubble caps, if possible.
Lower system pressure.
Add additional cooling to dry glycol and increase circulation rate.
Change out absorber to a small unit designed for a lower flow rate.
5.13.3.5 Absorber Tray Damage
Open inspection ports and/or manway and verify tray integrity. Repair or replace as necessary.
5.13.3.6 Breakdown or Contamination of Glycol
Have lean and rich glycol sample analyzed.
Note evidence of severe contamination, thermal or chemical decomposition. Clean system and/or recharge with fresh glycol as necessary.
5.13.3 Glycol Loss from the Contactor
5.13.4.1 Foaming
Major cause of foaming is contamination.
Remove source of contamination. Clean contactor if necessary, clean system if necessary, replace glycol if necessary.
Increase filter capacity and/or add carbon filtration.
Add antifoam compound (silicon emulsion type).
Adjust high pH to prevent emulsification (use acetic acid).
5.13.4.2 Plugged or Dirty Trays
Manually enter tower and clean.
Open inspection ports and clean with water jet or by hand.
Chemically clean.
5.13.4.3 Excessive Velocity
Decrease gas rate.
Increase absorber pressure.
5.13.4.4 Interrupted Liquid Seal on the Trays (Gas Surge)
If the contactor has a bypass valve, isolate the tower by opening the bypass valve and closing the gas inlet valve. Allow the glycol pump to run 5 minutes then while the glycol is circulating open the gas inlet valve and slowly close the gas bypass valve.
If contactor does not have a bypass valve, stop or greatly reduce the gas flow through the tower (shut wells, flare gas, alternate system, etc.). Allow the glycol to circulate 5 minutes then slowly turn the gas back through the tower.
If unable to stop or reduce the gas flow, increase the glycol circulation rate to the maximum possible for 2–5 minutes (flood trays in attempt to reestablish seal using liquid head pressure).
5.13.4.5 Cold Glycol (Cold Gas)
Increase gas temperature by increasing temperature of flowline heater or add flowline heater, if necessary.
5.13.4.6 Leaks
Perform a pressure test for external gas glycol heat exchanger to check or detect glycol leakage into dry gas stream.
Check drain header (if applicable) at all gauge columns, external float cages (LSLL, etc.).
5.13.4.7 Accumulation in Integral Scrubber
Check for communication between chimney tray and scrubber section.
Check bottom tray leakage.
Check glycol level control and dump valve operation (electric powered glycol pumps units).
5.13.5 Glycol Loss from the Reconcentrator
5.13.5.1 Leaks
Be sure all drain valves are closed.
Be sure gauge column seals are good.
Check heat tube integrity (glycol loss into fire tube or waste heat tube will produce heavy smoke from stack).
Check reconcentrator shell integrity (note glycol leakage from insulation, etc.)
Heat source flange leak (poor gasketing).
5.13.5.2 Bad Glycol Relief Valve
Check glycol relief valve, replace if necessary.
5.13.5.3 Exiting the Still Column
For plugged or fouled still column packing, clean or replace still column packing.
For saturated glycol (droplets blowing out still):
Check reconcentrator heat source. Insure heat is between 3500 and 400 0F.
Check for free liquid or misting liquid carryover into contactor tower.
Repair or replace separator control, if necessary.
Reduce slugging if possible. Add scrubber, if necessary.
Reduce glycol flow through the reflux condenser.
5.13.5.4 Vaporization
Check reconcentrator temperature (below 404 0F).
Check reflux temperature. Increase the glycol flow through the reflux condenser to lower the reflux temperature.
Check stripping gas flow rate.
Check for plugged or fouled glycol outlet from reconcentrator (downcomer or heat exchangers).
5.3.16 Glycol Loss From Glycol Hydrocarbon Separator
5.13.6.1 Improper Control Operation
Repair or replace level control.
Clean, repair, or replace dump valve.
Check gas velocity in separator and mist extractor.
5.13.6.2 Leaks
Check drain valve. Tighten, repair, or replace.
Check gauge columns, external float cages, and level control adapters.
Add antifoam compound to prevent loss through gas outlet.
5.13.6.3 Accumulation in Oil Bucket (Bucket & Weir)
Open vessel and clean glycol passage under oil bucket (horizontal vessels).
Adjust the weir.
5.13.7 Glycol Loss—Miscellaneous
5.13.7.1 Leaks
Check all flanges, unions and associated piping.
Check electric pump rod packing.
Check all drain valves (filter, heat exchanger, etc.).
Check pump bleed valves (and electric pump bypass).
Check external gas-glycol heat exchanger.
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Still Column Diameter Size
Diameter size is based on the required diameter at the base of the still, calculated by vapor and liquid loading conditions at that point. Vapor load consists of the water vapor (steam) and stripping gas flowing up through the still. Liquid load consists of the rich glycol stream and reflux flowing downward through the still column. The diameter required for the still is based on the glycol circulation rate (Figure 5-44).
Still Column Packing
1 to 3 theoretical trays (4 to 12 feet) is sufficient for most TEG stripping still requirements.
304 SS packing is normally used.
Amount of Stripping Gas
The amount of stripping gas required to reconcentrate the glycol to a high purity will range from 2 to 10 ft.3/gal TEG circulated (Figure 5-45).
5.7.15 Filters
Microfiber : Sized to remove 5 micron solids.
Activated Charcoal (Carbon) :
Used to remove chemical impurities. Sized for full flow with 10 gpm streams.
Sized for 10 to 25% side streams on large units.
Fig. 5-42 Still column with wet glycol entering above the ceramic saddle packing.
Fig. 5-43 Still column with wet glycol entering below the stainless steel pall rings.
Fig. 5-44 Determination of stripping still column diameter.
Fig. 5-45 Amount of stripping gas required to reconcentrate glycol to high purity.
5.8 Calculation Examples for Glycol Dehydration
Example 5-1
30 MMscf/d of a 0.65 sp gr natural gas enters a TEG contactor at 600 psia and 100°F. Outlet water content specification is 7 lb H2O/MMscf and the TEG circulation rate is 3 gal TEG/lb H2O. Estimate the contactor diameter and number of bubble cap trays or height of structured packing required to meet this requirement. Assume z = 0.92.
Gas mol. Wt. = 0.65 X 28.96 = 18.8
Gas density, ρg =P (MW)/ RTZ
= (600) (18.
/ (10.73) (560) (0.92) = 2.0 lb/ft3Liquid density “from table 4-4”, ρL = 1.119* (62.4) = 69.9 lb/ft3
From fig. 4-8. Water inlet (Win)= 90 lb/MMscf
Solutions Steps:
1. From Fig. 4-8 (McKetta and Wehe chart) is equivalent to a water content of 7 lb H2O/MMscf @ 600 psia)
2- Estimate required TEG concentration from Fig. 5-10, H2O Dewpoint = 24°F,
@ T = 100°F, lean TEG concentration ≈ 98.8 wt%
3. Estimate number of theoretical stages.
Calculate water removal efficiency
(Win −Wout) / Win = (90 −7) /90 = 0.922
From Fig. 5-37 (N = 1.5) at 3 gallon TEG/lb H2O and 99.0 wt% TEG
(Win – Wout)/Win = 0.885
From Fig. 5-38 (N = 2.0) at 3 gallon TEG/lb H2O and
99.0 wt% TEG
(Win – Wout)/Win = 0.925 use N = 2.0
Using equations 5-6, 5-7, and 5-8.
G = C [ρV (ρL - ρV)]0.5 Eq. 5-6
2 theoretical stages ≅ 8 bubble cap trays @ 24 inch tray Spacing Eq. 5-7
2 theoretical stages ≅10 ft of structured packing Eq. 5-8
“C” from table 5-3
4. Size the contactor Bubble caps, 24 inch tray spacing:
G = 576 [2.0 (69.9 −2.0]0.5 = 6712 lb /ft2.hr
Mass flow, m =
M = 62000 lb/hr
Area of flow (A) = m “mass flow” /G “mass velocity” = 62000 / 6712 = 9.2 ft2
Since, area ft2 = π D2 /4
D2 = 9.2 X 4 / 3.14 = 11.7
D = 3.4 ft.
Alternatively
using equation 5-5 ( drag coefficient CD =0.85 and dm = 150 micron)
d2 = 5040 [(T0ZQg)/P] [(ρg / ρL – ρg)(CD/dm)]0.5 Eq. 5-5
d2 = 5040 [(560 x 0.92 x 30 x 106)/600] [(2 / 67.9)(0.85/150)]0.5
minimum diameter D = 3.5 ft.
5. Size For Structured packing: ( “C” from table 5-3 “ use 1260 as an average”)
G = 1260 [2.0 (69.9 −2.0]0.5 = 14683 lb /(ft2.hr)
Mass flow, m = (30 X106) (0.65 X 28.96) / (379.5) (24) = 62000 lb/hr
Area A = m/G = 62000 / 14683 = 4.22 ft2
Area ft2 = π D2 /4
D2 = 4.22 X 4 / 3.14 = 5.38
D = 2.3 ft
Example 5-2
Determine reboiler duty for conditions in the previous example. Assume the rich TEG temperature entering the regenerator is 300°F and the reboiler temperature is 400°F.
Glycol Reboiler Duty: Basis 1 gal. TEG. ( assume 25% reflux ratio)
Solution:
Using equation 5-12, and 5-13
Qs = m Cp Δt Eq. 5-13
- (use 0.67 as an average for the range of 300-400 0F)
Qs = (9.3 lb/gal.) X (0.67 Btu/lb.°F) X (400°F − 300°F) = 623 Btu /gal.
Using eq. 5-14
Qv = (ΔHvap) (ΔW) Eq. 5-14
Qv = (970 Btu/lb H2O) X (1 lb H2O/ 3 gal. TEG) = 323 Btu / gal.
Using eq. 5-15
Qc = % Reflux Ratio X Qv /100 Eq. 5-15
Qc = (25) (Qv)/100 = 81 Btu /gal.
From eq. 5-12
Total Duty Including 10% Heat Loss:
Qr (total regeneration heat duty) = (623 + 323 + 81) (1.1) = 1130 Btu /gal.
Total Duty Based on 30 MMscfd of Gas:
Q = (1130 Btu/gal) (3 gal./lb) (30 MMscfd/24) ((90-7) lb/MMscf) = 350,000 Btu/hr
Example 5-3
Given: Gas Qg = 98 MMscfd at 0.67 SG saturated with water at 1000 psig and 100 0F
Dehydrate to = 7 lb/MMscf
Use triethylene glycol, No stripping gas is available
98.5% TEG concentration
CD (contactor) = 0.852
Z = 0.865
Determine:
1. Calculate contactor diameter
2. Determine glycol circulation rate.
3. Size the still column
Solution:
1. Calculate contactor diameter
d2 = 5040 [(T0ZQg)/P] [(ρg / ρL – ρg)(CD/dm)]0.5 Eq. 5-5
dm = 125 microns, T = 570 0R
P = 1015 psia, Qg = 98 MMscfd
ρL = 70 lb/ft.3
ρg = (0.67 X 1015)/(560 X 0.865) = 3.79 lb/ft3
d2 = 5040 [(560 X 0.865 X 98)/1015] [(3.79/70-3.79)(0.852/125)]0.5
d = 68.2 in.
Use 72.00 ID contactor (standard off-the-shelf)
2. Determine glycol circulation rate and reboiler duty
Wi = 63 lb/MMscf (from McKetta-Wehe) “saturated water content”
W0 = 7 lb/MMscf (spec)
ΔW = Wi - W0 = 63 - 7 = 56 lb/MMscf
ΔW/Wi = 56/63 = 0.889
Using n = 2 (i.e., 8 actual trays) and glycol purity of 98.5% read from Figure 5-38 the glycol circulation rate of about 2.7 gal TEG/lb H20. Use 3.0 gal/lb for design.
Using equation
L = (ΔW/Wi) Wi Qg/24 5-10
L = (3.0 gal/ Lb) (56 lb/MMscf)(98 MMscf/day)(day/24 hr)(hr/60min)
= 11.4 gpm TEG
= 862 Btu/gal (From table 5-4)
= (862 Btu/gal) (11.4 gal/min) (60 min/hr)
= 590,000 btu/hr
To allow for start-up heat loads, increase heat duty by 10% and then select a standard off-the-shelf fire tube.
Thus, select a 750 MMBtu/hr.
3. Design of still column:
Use 12-foot still column (standard packed arrangement)
dm = 125 micron
T = 300 0F = 760 0R
P = 1 psig = 15.7 psia
Qg = (10 scf/gal) (11 gal/min) (60 min/hr)(24 hr/day)
= 0.16 MMscfd
Z = 1.0
Since ρg = Gas density, lb/ft.3 = 2.7 (SP/TZ) or = ρg= 0.093 ((MW)P)/TZ lb/ft3 (Eq. 1-19)
ρg = 2.7 (0.62 X 15.7)/ (760 X 1.0 )
= 0.035 lb/ft.3
ρL = 62.4 lb/ft.3
CD = 14.2 (given)
d2 = 5040 [(760 X 1 X 0.16)/15.7] [(0.035/62.4-0.035)(14.2/125)]0.5
d = 17.7 in.
Use 18 inch OD x 12 feet long still.
5.9 Glycol Unit Operation
5.9.1 Start up
Prior to the initial start up of a new plant, the vessels and lines should be thoroughly washed out with water to remove debris and corrosion products that accumulated during construction. After the system has been cleaned, start up is accomplished in three phases:
Establish glycol circulation throughout the plant.
Apply heat to the reboiler and bring it up to operating temperature.
Open the wet gas stream to the contactor and begin dehydrating the gas.
In order to circulate glycol throughout the system, it will be necessary to pressurize the vessels in the system. Pressuring can be done with wet gas or dry gas. The contactor pressure should be raised to at least 150 psig (10 bars) and the flash tank pressure should be raised to at least 45 psig (3 bars). When the vessels have been pressured, start up procedure is:
Fill the reboiler and surge tank with fresh glycol solution Also add to the flash tank.
Pressure up the contactor column by very slowly opening the gas inlet valve.
Prime and start glycol pump.
When liquid appears at the bottom of the contactor put the bottom level controller in service so the glycol will flow to the flash tank.
Put the flash tank level controller in service when liquid appears in the bottom, so that liquid will flow to the stripper.
Keep surge tank level half full by adding glycol when needed.
When desired circulation rate is established, light the reboiler or put the heat source in service and slowly bring reboiler temperature up to 250°F (121°C). Leave temperature at 250°F (121°C) until all water has been boiled out of glycol.
5.9.2 Routing Operation
Routing operating checks include the following:
Check levels in each vessel and reset level controller as necessary.
Check the pressure drop across the filter and replace the elements as required.
Check the temperature of the lean glycol out of the glycol exchanger to see that the proper transfer rate is occurring in the exchanger.
Check the flow of glycol to the contactor and of stripping has to the reboiler.
Check the pressure of the flash tank to see if it is at proper level.
If water or air is used to cool the glycol prior to its entry into the contactor, check the glycol temperature in order to ensure that it is about 5°C to 7°C (10°F to 15°F) above the inlet gas temperature. Adjust the flow of air or water through the cooler as required.
Check the water content of the outlet gas to see that it is below the design limit.
5.9.3 Shut Down
This procedure is used to shut down a glycol plant:
Block gas to contactor column,
Shut off heat (leave pump running),
When unit cools to safe temperature, less than 200°F, shut off pump.
Drain glycol, if necessary.
5.10 Glycol Maintenance and Care
5.10.1 Preventive Maintenance
Scheduled preventative maintenance reduces glycol losses and operating problems such as foaming, system plugging, corrosion, pump failures.
It also minimizes system down time and maximizes system operation efficiency.
Five Steps to a Successful Preventive Maintenance Program
Record-Keeping
Accurate records can be used to determine the system efficiency and to pinpoint operating problems. Records of prior and existing conditions including dew points, glycol usage, and repairs help establish the system profile. Once the system profile is defined it becomes easier to identify unusual system characteristics that may indicate potential problems.
Mechanical Maintenance
Daily physical inspections are necessary to insure that the system is running properly. Any trouble encountered should be dealt with immediately, thus preventing the problem from escalating.
Glycol Care
Regular chemical analysis (every one or two months) of the glycol provides detailed information on the internal operation of the unit. Many process-related problems can be diagnosed well in advance of mechanical failure. Chemical problems can be diagnosed and corrective action taken before they become costly and detrimental to unit performance.
Corrosion Control
Corrosion is a frequent problem in glycol dehydration systems. If unchecked the damage can be extensive. All units should have provisions for corrosion control.
Communication
Lines of communication between field and technical assistance or office personnel are critical to the smooth operation of any system. Office personnel (production supervisors, facility engineers, purchasing agents) must be kept informed of daily operations and any problems that may arise.
Field personnel must be made aware of technical information that may improve their operations. Training for field operators allows the operator to better maintain the equipment.
5.10.1.1 Record-Keeping
Records necessary to establish a system profile include:
Design information including vessel specifications, equipment drawings, and P&IDs
Filter element or media replacement—type and frequency
Glycol usage—gallons/month
Chemical additives—type and amount
Gas production and flow rate charts—peak, average, and low periods
Outlet gas dew point/water content (lbs/Mscfd)
Mechanical inspections—type, magnitude, frequency, results
Records necessary to establish a system profile include:
Glycol analysis—format, frequency, recommendations, results
Corrosion coupon results—mills per year (MPY), frequency
Materials and labor relating to system repairs— operating costs
With the aforementioned information, a good system profile can be drawn of a specific system.
Updating these records will show any gradual changes in a unit’s system profile and may alert you to a potential or developing problem.
5.10.1.2 Mechanical Maintenance
The following things should be done so as to keep the unit operating properly and to prevent operational problems:
1. Insure that instruments and controls are in good working condition (thermometers and pressure gauges, etc.). Use a test thermometer on the reconcentrator to insure proper reconcentrator heat.
2. Insure glycol filter elements are changed according to average expected need basis:
Microfiber filters should be changed monthly.
Carbon filters should be changed monthly (small cartridge filters) to every six months (large bulk units). Glycol analysis helps determine the frequency.
An upset or sudden change in the operating conditions may foul the filters faster than the preventive maintenance program anticipates. Make sure filter differential pressure is below 15 psi.
3. Look for glycol leaks on and around the glycol skid. Most leaks can be stopped by tightening a union, valve stem packing, or pump rod packing. After the leak has been repaired, clean the affected area so it is easier to notice new leaks.
4. Check the glycol level, at least twice a day, and add glycol as necessary. Maintain a written report of glycol added. This allows operations to detect excessive losses of glycol and take corrective action faster.
5. Insure unit performance by taking a dew point measurement daily.
6. Clean the glycol strainers monthly to prevent accumulation of trash, which can cause the glycol pump to fail.
7. Check the glycol circulation rate daily. Any time the gas flow rate changes or when a drastic change in gas pressure or temperature is experienced, the glycol flow rate should be recalculated and the pumps set accordingly. On multiple pump installations, switch the pumps weekly, thus insuring pump operation when necessary.
8. On direct fired fire tubes, on weekly basis; sight down the flame direction, to insure it is not touching the fire tube for fire tube blisters or hot spots.
9. Cycle the main burner manually to be sure the fuel gas valve works and the pilot light stays lit. Check the fuel gas scrubber pot for fluid build-up that may hinder burner operation.
5.10.1.3 Glycol Care
Operating and corrosion problems occur when the circulating glycol gets dirty. Some contaminated glycol problems can be noticed easily and corrective action taken.
A small glycol sample should be taken daily from the surge tank or dry glycol suction header to the pump. Check closely for fine black particles settling out of the sample, which may be corrosion by-products and indicate an internal corrosion problem.
Uncontaminated glycol is colored like pure kitchen oil. Glycol gets darker and darker due to hydrocarbon contaminates, and/or corrosion products.
Smell the sample, If the smell is sweet and aromatic (similar to rotten bananas) it may be thermally decomposed. If the sample is viscous and black it is probably contaminated with hydrocarbon or well-treating chemicals. If the hydrocarbon contamination is great enough, the sample will separate into two liquids or interphase. Every one or two months send a sample of both the rich and lean glycol to a laboratory for complete analysis. This type of analysis will provide a detailed description of unit performance and glycol condition.
5.10.1.4 Corrosion Control
Corrosion is a major cause of premature equipment failure. Corrosion can occur over the entire system, inside and out. The two most common areas of severe corrosion are:
Still column reflux coil, and vent/fill connection on the surge tank.
This is due to a high concentration of water vapor in the top of the still and the ready availability of oxygen in the air at the vent/fill cap. Three types of corrosion are usually found in glycol systems either individually or in combination with one another are:
Oxidation
Is the corrosion process involved the presence of oxygen molecules. Some metal loss is incurred and the resultant corrosion product is a scale-like residue called oxide, or rust.
Oxidization is characterized by rough, irregular, shallow pitting of the metal scaled over the rust.
Sour Corrosion
Hydrogen sulfide (H2S) are often found in produced natural gas. Glycols are very reactive with sulfur compounds, such as H2S. The resulting materials tend to polymerize (form larger molecules) that form a “gunk” that is very corrosive.
Corrosion in the presence of acid gases is characterized by deep, jagged pitting.
Sweet Corrosion
Water is found in glycol as vapor, free condensed water, or entrained water in glycol. Carbon dioxide (CO2) when dissolved in water forms carbonic acid. Since most produced natural gases contain some CO2, the presence of carbonic acid in glycol systems is very common. The corrosion resulting from carbonic acid is known as sweet corrosion. Sweet corrosion is characterized by deep, round, smooth pitting. Sometimes the pitting will cover a broad area, disguising the depth of the pit.
5.10.1.4.1 Corrosion Monitoring and Control Programs
Monitoring and control programs should include system monitoring through:
Corrosion coupons
Glycol analysis (pH and iron)
Three steps in combating corrosion in glycol systems are:
Use an effective corrosion inhibitor in both the liquid and vapor phases.
Use corrosion resistant alloys (CRA) in construction.
Keep the unit clean to prevent acid formation due to contamination.
Cathodic protection has been attempted but met with little success.
It is impractical to attempt to eliminate corrosion. The rate of corrosion can be slowed to a point that is almost negligible. The maximum acceptable corrosion rate is 6 mils per year (MPY).
Corrosion inhibitors work in several ways.
The two most applicable to glycol units are:
pH buffers which include:
Alkanolamines
MEA (Monoethanolamine)
TEA
The pH buffers fight corrosion by stabilizing the pH near neutral, thereby reducing corrosive environment. Alkanolamines are regenerable as is the glycol and thus can be retained in the system for lengthy time periods. However, they are thermally degraded at normal reconcentrator operating temperatures and if used frequently may leave harmful residues within the system.
Plating Inhibitors
Tallow diamine, unlike the inorganic amines, is an organic amine. It is grouped with the plating inhibitors even though it does not actually plate out on the vessel walls. It flashes out of the glycol at high temperatures. As it vaporizes it contacts the vapor spaces of the reconcentrator and forms a tenacious film over the exposed metal. These inhibitors are strictly liquid phase protection. They will plate out on the vessel walls forming a protective barrier between a corrosive environment and the metal. This film will eventually wear away and must be replenished occasionally to continue protection.
Since the plating inhibitors are all alkalies, some degree of pH buffering will be effected.
The pH buffering will not be as great as through the use of amines.
True plating inhibitors include:
Borax
NaCap (Sodium mercaptobenzothiazole)
Dipotassium phosphate
5.10.1.5 Communication
Communication is the easiest portion of an effective maintenance program and yet it is the most overlooked. Communication can be between management and labor, engineer and foreman, operators on opposite shifts, and office and field personnel. Lack of communication is the single most contributing factor to glycol system failure. Failure to communicate can cause confusion and evolve into major problems.
5.11 Glycol Operation Considerations
Operating and corrosion problems usually occur when the circulating glycol gets dirty.
To achieve a long, trouble-free life from the glycol, it is necessary to recognize these problems and know how to prevent them.
Some of the major areas are:
Oxidation
Thermal decomposition
pH control
Salt contamination
Hydrocarbons
Sludge
Foaming
5.11.1 Oxidation
Sometimes glycol will oxidize in the presence of oxygen and form corrosive acids. Oxygen enters the system with the incoming gas through:
Unblanketed storage tanks and sumps
Pump packing glands
To prevent oxidation:
Bulk storage tanks should be gas blanketed.
Use oxidation inhibitors.
Normally, a 50/50 blend of MEA and 33.5% hydrazine is inserted into the glycol between the absorber and the reconcentrator. A metering pump should preferably be used to give continuous, uniform injection.
5.11.2 Thermal Decomposition
Excessive heat, a result of the following conditions, will decompose glycol and form corrosive products:
High reconcentrator temperature above the glycol decomposition level
High heat-flux rate, sometimes used by a design engineer to keep the heater cost low
Localized overheating, caused by deposits of salts or tarry products on the reconcentrator fire tubes or by poor flame direction on the fire tubes
5.11.3 pH Control
pH is a measure of the acidity or alkalinity of a fluid, based on a scale of 0 to 14. pH values from 0–7 indicate the fluid is acidic. pH values from 7–14 indicate the fluid is alkaline.
To obtain a true reading, glycol samples should be diluted 50-50 with distilled water before pH tests are run. The pH meter should be calibrated occasionally to keep it accurate.
New glycol has a neutral pH of approximately 7. With usage the pH decreases and the glycol becomes acidic and corrosive unless pH neutralizers or buffers are used. Equipment corrosion rate increases rapidly with a decrease in the glycol pH. Acid created by glycol oxidization, thermal
decomposition products, or acid gases picked up from the gas stream are the most troublesome of corrosive contaminants. In addition, a low pH accelerates the decomposition of glycol.
Ideally, the glycol pH should be held at a level between 7.0 and 7.5. A value above 8.5 tends to make glycol foam and emulsify. A value below 6.0 corresponds to system contamination, corrosion, and/or oxidation.
Borax, ethanolamines (usually triethanolamine) or other alkaline neutralizers are used to control the pH. These neutralizers should be added slowly and continuously for the best results.
An overdose will usually precipitate a suspension of black sludge in the glycol. The sludge could settle and plug the glycol flow in any part of the circulating system.
Frequent filter element changes should be made while pH neutralizers are added.
The amount of neutralizer to be added and the frequency will vary from location to location. Normally, 1/4 lb of triethanolamine (TEA) per 100 gallons of glycol is sufficient to raise the pH level to a safe range.
When the glycol pH is extremely low, the required amount of neutralizer can be determined by titration. For best results, the lean rather than the rich glycol should be treated. It takes time for the neutralizer to mix thoroughly with all the glycol in the system. Several days are required before the pH is raised to a safe level. Each time that neutralizer is added, the pH of the glycol should be measured several times.
5.11.4 Salt Contamination and Deposits
In areas where large quantities of brine are produced, some salt contamination will occur. Salt deposits accelerate equipment corrosion. It also reduces heat transfer in the fire tubes. It alters specific gravity readings when a hydrometer is used to measure glycol water concentration. It cannot be removed with normal regeneration. A scrubber installed upstream of the glycol plant should be used to prevent salt carry-over from produced free water. The removal of salt from the glycol solution is then necessary.
5.11.5 Hydrocarbons
TEG will typically absorb about 1 scf of sweet natural gas per gallon of glycol at 1000 psia and 100°F. Solubilities will be considerably higher if the gas contains significant amounts of CO2 and H2S. Heavier paraffin hydrocarbons are essentially insoluble in TEG. Aromatic hydrocarbons, however, are very soluble in TEG, and significant amounts of aromatic hydrocarbons may be absorbed in the TEG at contactor conditions. This may present an environmental or safety hazard when they are discharged from the top of the regenerator.
Vapor-liquid equilibrium constants (K-values) for benzene, toluene, ethylbenzene, and o-xylene in TEG solutions, indicates that at typical contactor conditions approximately 10-30% of the aromatics in the gas stream may be absorbed in the TEG solution.
Aromatic absorption increases with increasing pressure and decreasing temperature. Aromatic absorption is directly related to TEG circulation rate. Higher circulation rates result in increased absorption. Aromatic absorption is essentially independent of the number of contacts in the absorber so one method of minimizing aromatic absorption is to use taller contactors and minimize TEG circulation rates.
Most of the aromatic components will be stripped from the TEG solution in the regenerator.
Flash tank sizing should be sufficient to degas the glycol solution and skim entrained liquid hydrocarbons, if necessary. A minimum retention time of 3-5 minutes is required for degassing. If liquid hydrocarbons are to be removed as well, retention times of 20-30 minutes may be required for adequate separation. Flash tank pressures are typically less than 75 psia.
Regenerator sizing requires establishing the reboiler duty and, when high TEG concentrations are required, providing sufficient stripping gas.
Liquid hydrocarbons, a result of carry-over with the incoming gas or condensation in the contactor, affects glycol by foaming, degradation, and losses. It must be removed with glycol/gas/condensate separator, or hydrocarbon liquid skimmer, or activated carbon beds
5.11.6 Sludge
Solid particles and tarry hydrocarbons (sludge) are suspended in the circulating glycol, and with time will settle out (It looks like asphalt or paraffin sludge). This results in the formation of black, sticky, abrasive gum that can cause trouble in pumps, valves and other equipment, usually when the glycol pH is low.
5.11.7 Foaming and defoamers
Excessive turbulence and high liquid-to-vapor contacting velocities usually cause the glycol to foam (this condition can be caused by mechanical or chemical problems).
The best way to prevent foaming is proper care of the glycol, such as:
Effective gas cleaning ahead of the glycol system
Good filtration of the circulating solution
Defoamers serve only as a temporary control until the conditions generating foam can be identified and removed. Success depends on when and how it is added. Most are inactivated within a few hours under high temperature and pressure, and thus their effectiveness is dissipated by the heat of the glycol solution. Thus, defoamers should be added continuously, a drop at a time, for best results.
The chemical feed pumps should meter the defoamer accurately, improve dispersion into the glycol, and may be activated automatically by differential pressure across the contactor.
5.12 Analysis and Control of Glycol
Analysis of glycol is essential to good plant operation. It helps pinpoint high glycol losses, foaming, corrosion, and other operating problems. Analyses enable operations personnel to evaluate plant performance and make operating changes to obtain maximum drying efficiency.
5.12.1 Visual Inspection
A glycol sample should first be visually inspected to identity some of the contaminants.
A finely divided black precipitate may indicate the presence of iron corrosion products.
A black, viscous solution may contain heavy, tarry hydrocarbons.
The characteristic odor of decomposed glycol (a sweet aromatic odor) usually indicates thermal degradation. A two-phase liquid sample usually indicates the glycol is heavily contaminated with hydrocarbons. The visual conclusion should next be supported by a chemical analysis.
5.12.2 Chemical Analysis
A complete glycol analysis of lean and rich samples, when properly interpreted, can provide a detailed picture of the workings of the dehydration unit and its process.
Glycol analysis should include tests to determine the following (table 5-5):
Test Lean Glycol Rich Glycol Allowable Range Ideal
pH (50/50) 6 to 8 7 to 7.5
Hydrocarbon
(% wt.) 0.1%
Water content
(% wt.) 2% lean - 6% rich
TSS (% wt.) 0.01%
Residue (% wt.) 4% 2%
Chlorides (mg/l) 1500 1000
Iron (mg/l) 50 35
Foam character:
Height (ml) 20 to 30 ml
Stability (sec) 15 to 5 sec
Specific gravity 1.118 to 1.126
Table 5-5.Glycol analysis.
5.12.3 Chemical Analysis Interpretation
5.12.3.1 pH
A pH below 6 generally corresponds with system contamination, corrosion, and/or oxidation.
Below 5.5 autoxidation occurs. Where Chemical decomposition of the glycol occurs within itself, and glycol has the tendency to continue to drop without outside influences.
The causes of low pH are:
Acid gases in the gas stream
Organic acids due to oxidation or thermal degradation
Excessive chlorides (salt) in the glycol
Well-treating chemicals entrained in the gas stream
Thermal decomposition of entrained liquid hydrocarbons in the gas stream and glycol
Oxidation of the glycol due to improper storage
While the causes of high pH are:
Contamination from well-treating chemicals entrained in the gas stream
Overdose of neutralizer added to a system for low pH
Foaming tendencies can result from high pH, due to stabilized glycol-hydrocarbon emulsions.
Sludge and residue build-up can result from both high and low pH.
Glycol pH should be checked periodically and kept on the alkaline side by neutralizing the acidic compounds with alkaline chemicals, such as monoethanolamine (MEA). A pH of about 7.3 is considered a safe level. Raising the pH above 8 to 8.5 is not desirable because of the tendency for an alkaline glycol solution to foam and emulsify more easily.
5.12.3.2 Sludge
Sludge may become abrasive and cause premature pump and valve failure. It may deposit in trays and downcomers, still column packing, and heat exchangers, which cause system plugging.
5.12.3.3 Hydrocarbons content
Enter the glycol stream as a result of inlet separator carry-over or as condensation due to temperature variations. Compressor lube oils and other extraneous organic chemicals such as pipeline corrosion inhibitors, are often stripped out of natural gas as it passes through the contact tower. Oils and organic residues can cause glycol/water emulsions and suspensions, which contribute to foaming, which will results in excessive high glycol carryover from the contactor, and the contaminants may cause plugging in the contactor, still column, and heat exchangers
The light hydrocarbons are usually separated from the glycol stream with an adequately sized glycol/hydrocarbon separator, while heavy hydrocarbons that are referred to as soluble hydrocarbons because they bond with the glycol are usually filtered out with activated carbon.
Light end hydrocarbons (insoluble) are allowable up to 1% by volume.
Soluble hydrocarbons are only acceptable to 0.1 % by weight, since they are primarily responsible for foaming, sludge and residue build up, low pH, loss of hygroscopicity, and glycol decomposition.
Hydrocarbons are often get into glycol in;
Condensation which is caused when glycol enters the contactor colder than the incoming gas. This problem can be eliminated by maintaining a temperature for the entering lean glycol of 10-15°F (5-7°C) warmer than the incoming gas.
Carryover of hydrocarbon contaminants from the inlet separator or gas
5.12.3.4 Water Content
Water content is defined as the quantity of water in the glycol. The difference between the lean sample and rich sample measures the degree of loading in the contactor. It indicates regeneration efficiency. Glycol purity should be at least 98% in the lean stream and at least 94% in the rich. These concentrations will produce the desired dew points in systems that are operating properly. For lower dew points the glycol purity must be increased (or water content decreased). High water content of the lean sample generally indicates low reconcentrator heat or one of the following reasons:
Excessive glycol circulation
Undersized equipment
Carry-over from the separator
Vapor communication from reconcentrator to surge
A leak in the glycol/glycol heat exchanger
Over-refluxing in the still column
Hot inlet gas temperature
High water content in the rich sample usually indicates a low glycol circulation rate or:
Carry-over from the separator
Poor reconcentration
Heat exchanger communication
Undersized equipment
Hot inlet gas temperature
Check values for hydrocarbon, chlorides, iron, and foaming to help pinpoint the problem.
5.12.3.5 Suspended Solids
Considered to be those solids and tarry hydrocarbons that remain suspended within the glycol solution down to 0.45 micron in size. They are result of poor inlet separation, corrosion, and thermal degradation of the glycol. Values greater than 0.01% by weight indicate poor sock/microfiber filtration. Most filters are sized to remove particles to a size of 5 microns.
Particles larger than this in excessive amounts may serve to stabilize foaming tendencies in glycol. When the glycol is allowed to maintain a large concentration of suspended solids, a silty residue is likely to form along vessel walls causing plugging of the contractor trays, heat exchangers, still column, and reconcentrator. Suspended solids are likely (common with low glycol pH).
Problems resulting from a high solids content include:
Increased pump wear from abrasion
Accelerated corrosion and erosion
Increased fouling of fire tube .
Increased glycol loss due to foaming
Increased plugging problems.
5.12.3.6 Residue
The value for residue is a function of system contamination. The glycol sample is distilled, removing all light end hydrocarbons, water, and glycol. Residue represents the remaining contamination, which is comprised of total solids (suspended and residual), salt, and heavy hydrocarbons.
Value for residue is best kept below 2% by weight, however some systems may operate reasonably well at values from 2% to 4%. Units with Glycol containing greater than 4% are prime candidates for failure and should be cleaned immediately.
5.12.3.7 Chlorides
Chloride values indicate the quantity of inorganic chlorides (salts) found in the glycol sample.
As the concentration of chlorides (as NaCI or CaCI) in glycol increases, its solubility decreases.
When heat is added to the glycol solution, the salt begins to form crystals which:
Fall out of the glycol solution
Accumulate on the heat source and can lead to premature heat tube failure
May be swept by the glycol into other areas of the system
Potential problems with excessive chlorides include system plugging, low pH, glycol pump damage, foaming, and loss of hygroscopicity due to rapid glycol decomposition.
Removal of chlorides in high concentrations requires vacuum distillation of the glycol.
Concentrations greater than 1000 ppm will stabilize foaming tendencies, may lead to excessive glycol loss, and may affect glycol pH.
Precipitation of salts from the glycol will begin at approximately 1200 to 1500 ppm, however, the crystals formed are extremely small and rarely troublesome. At concentrations above 2200 ppm, precipitation occurs readily and system failure is a possibility.
Filtration removes large salt crystals, but most of the damage associated with salt will have already occurred prior to the development of crystals large enough to filter.
5.12.3.8 Iron
Iron found in glycol samples can indicate possible corrosion, and/or produced water carry-over.
Iron in excess of 50 ppm generally indicates corrosion. Whether it be in the glycol unit, upstream in the production equipment or downhole in the well string is difficult to determine.
Comparing values for iron content in several points downstream the glycol unit helps to establish the location of suspected corrosion. Corrosion by-products will consist of soluble iron and fine, gritty particulate in systems where oxygen is available. In systems where no oxygen is present, corrosion by-products will include sulfides in addition to the iron.
5.12.3.9 Foaming
More glycol is lost through foaming than any other cause. Foaming is not easily detected without chemical analysis; gradual low-volume glycol loss often goes overlooked. It is usually a result of contamination. Primary contaminants that cause foaming are hydrocarbons (from separator carryover), suspended solids, chlorides, compressor lube oil, well treating chemicals, and iron.
Water content affect foaming tendencies by inducing emulsification of contaminants, particularly hydrocarbons. Carbon filtration is the most effective means of controlling foam. Silicone emulsion–type foam inhibitor is used, but they treat the symptom, not the cause and thus are temporary solutions. Addressing the source of the contamination causing the foam is the only long-term solution.
The foam test consists of bubbling dry air at a rate of 6 liters/min through a graduated cylinder container of 200 mm of the glycol sample until the foam stabilizes at its maximum height.
Volume for both the liquid and the foam is reported as a single value. The original 200 ml is then subtracted. The remaining value is recorded as height and represents the ease at which the solution will foam. Once the maximum foam height is recorded, the dry air is removed from the sample and the time it takes for the foam to break from its maximum volume to a clear surface on the glycol sample is recorded in seconds. This time represents the tendency of the foam and is known as stability.
There are no concrete values given for acceptable foam height and stability. Foam with very low height and moderate stability will result in little glycol loss as will a foam with moderate height and very low stability. Thus, the acceptable range for foam test results are:
Height/ml: 20 to 30 ml
Stability/sec: 15 to 5 sec
For example, a sample with a height of 25 ml and a stability of 10 sec is acceptable, while a sample with 30 ml height and 15 sec stability would have a high foaming tendency and could result in glycol losses.
5.12.3.10 Glycol Weight Percentage
This refers to the amount of glycol in the glycol solution. Lean glycol should contain about 98.5 to 99.9% a glycol. Rich glycol content varies from about 93-96% glycol.
5.12.3.11 Specific Gravity
Specific gravity is used to determine the purity of glycol. A specific gravity of 1.126 to 1.128 at 60 0F indicates a 99% TEG (technical grade).
A specific gravity of 1.124 to 1.126 indicates 97% (industrial grade).
With glycol extracted from an operating dehydration unit, the lean sample should have a specific gravity of 1.1189 to 1.121.
This variance allows for acceptable amounts of system contamination.
Low specific gravity would indicate one or more of the following:
TEG containing excessive amounts of EG and/or DEG (poor quality replacement glycol)
Excessive water in sample
Excessive hydrocarbons in sample
A high specific gravity indicates one or more of the following:
The system is contaminated with excessive amounts of solids or any additives with a greater density than glycol
Thermal degradation of the glycol
Oxidation or chemical degradation of the glycol
5.12.3.12 Glycol Composition
The composition of glycol indicates its quality.
Values are given to the component glycols (EG, DEG, TEG, TTEG) contained within the glycol sample solution. Industrial grade (97%) TEG or better is required for best glycol system results.
In addition to 97% TEG, the glycol solution may contain, in various concentrations, up to 1% EG and 3% DEG, but not to exceed a combined total of 3%.
Glycol degradation will often be reflected by changes in the glycol composition and reduction in pH. Thermal degradation is most common and is characterized by excessive values of EG, DEG, and occasionally the presence of TTEG. The thermal degradation is characterized by:
The glycol pH will be low.
The glycol sample will be dark and have an aromatic smell (ripe bananas).
Chemical degradation is brought about by oxidation and acidic contaminants and is characterized by:
Excessive values for EG and DEG but no TTEG will be present
Low pH
Glycol may not appear to be too dirty
Autoxidation is a form of continuing chemical degradation.
5.13 Troubleshooting
5.13.1 General Considerations
The most obvious indication of a unit malfunction is high water content (dew point) of the outlet stream. High water content is brought about by:
Insufficient glycol circulation
Insufficient reconcentration of the glycol
These problems can be caused by a variety of contributing factors such as mechanical causes or existing operating conditions for which the equipment was not designed
5.13.2 Main approach to troubleshooting:
Timeframe
Determine the approximate date/time at which the problem became apparent.
List Changes
Inventory any changes (things that happened differently than usual). Look for what is different.
Production changes
Operational changes
Maintenance and Repairs
Weather
Investigate by process of elimination reduce the list of changes to determine the factor or factors that manifest the problem.
5.13.3 High Dew Points
5.13.3.1 Insufficient Glycol Circulation
If there is insufficient glycol circulation, check heat exchangers and glycol piping for restrictions or plugging.
On an electric driven piston pump:
Check flow rate indicator (if present) to insure proper glycol circulation. If flow rate indicator is not present, verify circulation rate by closing the glycol discharge valve from the contactor and timing the fill rate in the gauge column.
Check high-pressure dry-glycol bypass valve. Close if necessary.
Check pump prime by shutting pump down, closing the discharge valve, opening the bypass valve and restarting the pump. Allow to run briefly under no load through the bypass line to remove any trapped gas in the pump.
On glycol-gas powered pumps:
Close dry discharge valve. If pump continues to run, open dry discharge bleed valve and allow running a few strokes. Once all gas is purged from put, close the bleed valve. If pump continues to run, discontinue use and send in for repair. If pump will not prime, but continues to run gas through the dry discharge bleed valve then:
Check pump suction strainer for plugging.
Check glycol level in surge tank.
5.13.3.2 Insufficient Reconcentration
Verify reconcentration temperature with test thermometer (3500 to 400 0F). Raise temperature if necessary.
Check glycol-to-glycol heat exchanger for leakage of wet glycol into the dry glycol stream.
Check stripping gas if applicable. Be sure stripping gas is in service at the proper rate.
Check for communication between the reconcentration vapor space and the surge tank vapor space.
5.13.3.3 Operating Conditions Different from Design
Check operation of upstream separators and scrubbers. Be sure not to overload system.
Increase absorber pressure. This may require installation of a back pressure valve.
Reduce gas temperature, if possible.
Increase circulation rate, if possible.
Increase reconcentrator temperature, if possible.
5.13.3.4 Low Flow Rate
Blank off a portion of the bubble caps, if possible.
Lower system pressure.
Add additional cooling to dry glycol and increase circulation rate.
Change out absorber to a small unit designed for a lower flow rate.
5.13.3.5 Absorber Tray Damage
Open inspection ports and/or manway and verify tray integrity. Repair or replace as necessary.
5.13.3.6 Breakdown or Contamination of Glycol
Have lean and rich glycol sample analyzed.
Note evidence of severe contamination, thermal or chemical decomposition. Clean system and/or recharge with fresh glycol as necessary.
5.13.3 Glycol Loss from the Contactor
5.13.4.1 Foaming
Major cause of foaming is contamination.
Remove source of contamination. Clean contactor if necessary, clean system if necessary, replace glycol if necessary.
Increase filter capacity and/or add carbon filtration.
Add antifoam compound (silicon emulsion type).
Adjust high pH to prevent emulsification (use acetic acid).
5.13.4.2 Plugged or Dirty Trays
Manually enter tower and clean.
Open inspection ports and clean with water jet or by hand.
Chemically clean.
5.13.4.3 Excessive Velocity
Decrease gas rate.
Increase absorber pressure.
5.13.4.4 Interrupted Liquid Seal on the Trays (Gas Surge)
If the contactor has a bypass valve, isolate the tower by opening the bypass valve and closing the gas inlet valve. Allow the glycol pump to run 5 minutes then while the glycol is circulating open the gas inlet valve and slowly close the gas bypass valve.
If contactor does not have a bypass valve, stop or greatly reduce the gas flow through the tower (shut wells, flare gas, alternate system, etc.). Allow the glycol to circulate 5 minutes then slowly turn the gas back through the tower.
If unable to stop or reduce the gas flow, increase the glycol circulation rate to the maximum possible for 2–5 minutes (flood trays in attempt to reestablish seal using liquid head pressure).
5.13.4.5 Cold Glycol (Cold Gas)
Increase gas temperature by increasing temperature of flowline heater or add flowline heater, if necessary.
5.13.4.6 Leaks
Perform a pressure test for external gas glycol heat exchanger to check or detect glycol leakage into dry gas stream.
Check drain header (if applicable) at all gauge columns, external float cages (LSLL, etc.).
5.13.4.7 Accumulation in Integral Scrubber
Check for communication between chimney tray and scrubber section.
Check bottom tray leakage.
Check glycol level control and dump valve operation (electric powered glycol pumps units).
5.13.5 Glycol Loss from the Reconcentrator
5.13.5.1 Leaks
Be sure all drain valves are closed.
Be sure gauge column seals are good.
Check heat tube integrity (glycol loss into fire tube or waste heat tube will produce heavy smoke from stack).
Check reconcentrator shell integrity (note glycol leakage from insulation, etc.)
Heat source flange leak (poor gasketing).
5.13.5.2 Bad Glycol Relief Valve
Check glycol relief valve, replace if necessary.
5.13.5.3 Exiting the Still Column
For plugged or fouled still column packing, clean or replace still column packing.
For saturated glycol (droplets blowing out still):
Check reconcentrator heat source. Insure heat is between 3500 and 400 0F.
Check for free liquid or misting liquid carryover into contactor tower.
Repair or replace separator control, if necessary.
Reduce slugging if possible. Add scrubber, if necessary.
Reduce glycol flow through the reflux condenser.
5.13.5.4 Vaporization
Check reconcentrator temperature (below 404 0F).
Check reflux temperature. Increase the glycol flow through the reflux condenser to lower the reflux temperature.
Check stripping gas flow rate.
Check for plugged or fouled glycol outlet from reconcentrator (downcomer or heat exchangers).
5.3.16 Glycol Loss From Glycol Hydrocarbon Separator
5.13.6.1 Improper Control Operation
Repair or replace level control.
Clean, repair, or replace dump valve.
Check gas velocity in separator and mist extractor.
5.13.6.2 Leaks
Check drain valve. Tighten, repair, or replace.
Check gauge columns, external float cages, and level control adapters.
Add antifoam compound to prevent loss through gas outlet.
5.13.6.3 Accumulation in Oil Bucket (Bucket & Weir)
Open vessel and clean glycol passage under oil bucket (horizontal vessels).
Adjust the weir.
5.13.7 Glycol Loss—Miscellaneous
5.13.7.1 Leaks
Check all flanges, unions and associated piping.
Check electric pump rod packing.
Check all drain valves (filter, heat exchanger, etc.).
Check pump bleed valves (and electric pump bypass).
Check external gas-glycol heat exchanger.
Fundamentals of Oil and Gas Processing
Basics of Gas Field Processing
Basics of Corrosion in Oil and Gas Industry
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Basics of Gas Field Processing
Basics of Corrosion in Oil and Gas Industry
https://oilprocessing.net
https://oilprocessing.net/oil/index.php
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Re: Basics of Gas Field Processing Book "Full text"
Gas Dehydration - Chapter 5 - Part 3
5.14 Glycol System Cleaning
Chemicals are frequently needed to clean the glycol system. If chemical cleaning is done properly, it can be quite beneficial to plant operations. While if done poorly, it can be quite costly and create long-lasting problems.
The most effective type of cleaner is a very heavy-duty alkaline solution. To provide optimum cleaning, the concentration, temperature, and pumping rate of the solution must be carefully controlled and an experienced, reputable vendor employed. A cascading technique can be used to save on the cost of cleaning chemicals.
Cleaning Techniques to Avoid
Steam cleaning is not effective and can be damaging and dangerous. It tends to harden the deposits in the system, making them almost impossible to remove.
The use of cold or hot water, with or without high detergent soaps, will do little good in cleaning the system. High-detergent soaps can create a serious problem by leaving trace quantities of soap after the cleaning job. Soap traces left in the system can make glycol foam for a long time.
Acid cleaning is good for removing inorganic deposits. Since most deposits in the glycol system are organic, acid cleaning is not very effective. It can easily create additional problems in the glycol system after the cleaning job.
5.15 Eliminating Operating Problems
Most operating problems are caused by mechanical failure. It is important to keep equipment in good working order. Following operating and maintenance suggestions helps provide a trouble-free operation.
5.15.1 Inlet Scrubber/Microfiber Filter Separator
The cleaner the inlet gas entering the absorber, the fewer operating problems there will be.
Potential problems if there was not an inlet scrubber of filter separator are:
Liquid water carryover, results in one or more of the following:
Dilutes the glycol
Lowers the absorber efficiency
Requires a greater glycol circulation rate
Increases the vapor-liquid load on the still column
Floods the still column
Vastly increases the reboiler heat load and fuel gas requirements
These problems also cause:
High glycol losses
Wet sales gases
If the water contained salt and solids, they would be deposited in the reboiler to foul the heating surfaces and possibly cause them to burn out.
The scrubber also prevents salt or other solids from entering the glycol system, where they could be deposited in the reconcentrator to:
Foul the heating surfaces
Burn out as hot spots
If liquid hydrocarbons were present:
They would pass onto the still column and reboiler
Lighter fractions would pass overhead as vapor and could create a fire hazard
Heavy fractions would collect on the glycol surface in the storage tank and could overflow the system.
Flashing of the hydrocarbon vapor can flood the still column and vastly increase the reboiler heat load and result in glycol losses.
Well corrosion control program should be planned and coordinated to prevent glycol contamination.
Scrubber or filter separator may be an integral part of the absorber or preferably a separate vessel. Vessel should be large enough to remove all solids and free liquids to keep these impurities from getting into the glycol system. Vessel should be regularly inspected to prevent any malfunction. Liquid dump line should be protected from freezing during cold weather.
Separator should be located close to the absorber so the gas does not condense more liquids before it enters the absorber.
Sometimes an efficient mist extractor, which removes all contaminants over one micron is needed between the inlet separator and the glycol plant to clean the incoming gas; this is particularly useful when paraffin and other impurities are present in a fine vapor form
When gas is compressed prior to dehydration, a coalescing type of scrubber (microfiber filter separator) placed ahead of the absorber insures removal of compressor oil in vapor form. Compressor oil and distillate can coat the tower packing either in the absorber or still column and decrease its effectiveness.
5.15.2 Absorber
This vessel contains valve or bubble cap trays or packing to give good gas-liquid contact.
Cleanliness is very important to prevent high sales gas dew points caused by foaming and/or poor gas-liquid contact. Plugged trays or packing could also increase glycol losses.
Unit startup considerations are as follows:
The pressure on the absorber should be slowly brought up to the operating range and then the glycol should be circulated to get a liquid level on all trays.
Next, the gas rate going to the absorber should be slowly increased until the operating level is reached.
If the gas enters the absorber before the trays are sealed with liquid, it will pass through the downcomers and bubble caps. When this condition occurs and the glycol is pumped into the absorber, the liquid has difficulty in sealing the downcomers, and the liquid will be carried out with the gas stream instead of flowing to the bottom of the absorber.
Gas flow rate should be increased slowly when changing from a low to a high flow rate.
Rapid surges of gas through the absorber may cause; Sufficient pressure drop through the trays to break the liquid seals, and/or Glycol to be lifted off the trays, which will flood the mist extractor and increase glycol losses
Unit shutdown considerations are as follows:
First, the fuel to the reboiler should be shut down.
Then the circulating pump should be run until the reboiler temperature is lowered to approximately 2000F (94 0C).
This precaution will prevent glycol decomposition caused by overheating.
The unit can then be shut down by slowly reducing the gas flow to prevent any unnecessary shocks on the absorber and piping.
The unit should be depressurized slowly to prevent a loss of glycol.
The dehydrator should always be depressurized from the downstream (gas outlet) side of the absorber.
A dehydrator installed on the discharge side of a compressor should be equipped with a check valve in the inlet line, located as close as possible to the absorber. Experience has shown that some glycol is sucked back into this line when a compressor backfires or is shut down.
Internal absorber damage to the trays and mesh pad may also occur with a compressor failure.
The installation of the check valve usually eliminates this problem.
All compressors taking gas from or feeding gas to a dehydrator should have pulsation dampeners. The absence of this safety device may cause fatigue failure of instruments, trays, coils, mesh pads and other parts of the dehydrator.
The glycol dump valve and level controller should be set for throttling action to give an even flow of glycol to the regenerator. This will prevent slugs, which could flood the stripper and cause excessive glycol losses.
The absorber must be vertical to insure the proper flow of glycol in the vessel and adequate contact of the glycol and gas. Inspection ports at the trays can be very useful when inspecting or cleaning the vessel.
If dry gas from a glycol unit is used for gas lift, care must be used in both sizing and operating the unit because of the unsteady gas rate required in this service. Control valves or backpressure valves are used to prevent a sudden overloading of the absorber which can break the downcomer seals in a tray type of vessel and cause excessive loss of glycol in the sales gas.
Absorbers sometimes need to be insulated when excessive condensation of light hydrocarbons collect on the vessel walls. This often occurs when dehydrating rich, warm gases in cold climates.
These very light hydrocarbons can cause tray flooding in the absorber and excessive glycol losses from the regenerator.
The mist extractor should receive special attention because glycol entrainment and well-crawling are difficult to effectively control.
The type and thickness of the mesh pad should be carefully studied to minimize glycol losses.
Care should also be taken after installation to avoid mesh pad damage. The maximum pressure drop through the contractor to avoid damage to the mesh pad is approximately 15 psi.
5.15.3 Glycol-Gas Heat Exchanger
Most units are supplied with a glycol-gas heat exchanger that uses the gas leaving the absorber to cool the lean glycol entering the absorber. This exchanger may be a coil in the top of the absorber or an external one. A water-cooled exchanger may be used when heating of the gas must be avoided. This exchanger may accumulate deposits, such as salt, solids, coke or gum which foul the heat exchanger surface, reduce the heat transfer rate, and increase the lean glycol temperature. All of the above increase glycol losses and make dehydration difficult. The vessel should be inspected regularly and cleaned when needed.
5.15.4 Lean Glycol Storage Tank or Accumulator
Normally this vessel contains a glycol heat exchanger coil which cools the lean glycol coming from the reboiler and preheats the rich glycol going to the stripper
The lean glycol is also cooled by radiation from the shell of the storage tank.
This accumulator should normally be insulated. Water cooling can also be used to help control the lean glycol temperature.
On conventional regenerators without stripping gas:
Accumulator must be vented to prevent trapping gas
Vapors, trapped in the storage tank, could cause the pump to vapor lock
A connection is usually provided in the top of the storage tank for venting
Vent line should be piped away from the process equipment and should not be connected to the stripper vent because this could cause steam to dilute the concentrated glycol
Some units are equipped to provide a dry gas blanket (no oxygen or air) in the storage tank. Blanket gas is normally piped to the regular vent connection on top of the storage tank. If blanket gas is used, it is commonly taken from the fuel gas line. Only a very slight flow of gas is required to prevent steam generated in the reboiler from contaminating the regenerated glycol.
The vessel should be inspected occasionally to see that sludge deposits and heavy hydrocarbons are not collecting in the bottom of the vessel. The heat exchanger coil should be kept clean so proper heat transfer can be made. This also prevents corrosion. If the heat exchanger develops a leak, the water rich glycol could dilute the lean glycol.
Glycol level in the storage tank should be checked and a level in the gauge glass should always be maintained. Glycol should be added as the level is pumped down. Records of the amount of glycol added should be maintained. Make certain the storage tank is not overfilled.
5.15.5 Stripper or Still Column
The stripper, or still column, is generally a packed column located on top of the reboiler to separate the water and glycol by fractional distillation. Packing is usually a ceramic saddle but 304 stainless steel pall rings can be used. A standard stripper usually has a finned atmospheric condenser in the top to cool the steam vapors and recover the entrained glycol.
Atmospheric condenser depends upon air circulation to cool the hot vapors. Increased glycol losses can occur on extremely hot days when insufficient cooling in the condenser causes poor condensation. High glycol losses can also occur on extremely cold, windy days when excessive condensation (water and glycol) overloads the reboiler. Excess liquids percolate out the stripper vent.
If stripping gas is used, an internal reflux coil is normally provided to cool the vapors to prevent excessive glycol losses. This is due to a larger mass of vapor leaving the stripper which will carry glycol. Adequate reflux is provided by passing the cool, rich glycol from the absorber through the condenser coil in the stripper. If properly adjusted, it can provide uniform condensation throughout the year.
Manual/automatic valve in the piping is furnished to bypass the reflux coil.
Under normal circumstances this valve will be closed and the total flow will be through the reflux coil. In cold weather operation, with extreme low ambient temperatures, this could produce too much reflux and the regenerator could become overloaded. Therefore, a portion or all of the rich glycol solution should bypass the reflux coil. This is accomplished by opening the manual/automatic valve until the reboiler can hold the temperature. This lowers the amount of reflux produced by the coil and reduces the load on the reboiler.
Sometimes a leak can develop in the cool glycol reflux coil in the top of the stripper. When this happens, excess glycol can flood the tower packing in the still column, upset the distillation operation, and increase glycol losses.
Broken, powdered packing can cause solution foaming in the stripper and increase glycol losses.
Packing is usually broken by excessive bed movement which is caused when hydrocarbons flash in the reboiler. Careless handling when installing the packing can also cause powdering.
As particles break down, the pressure drop through the stripper increases. This restricts the flow of vapor and liquid and causes the glycol to percolate out the top of the stripper.
Dirty packing, caused by sludge deposits of salt or tarry hydrocarbons, will also cause solution foaming in the stripper and increase glycol losses. Packing should be cleaned or replaced when plugging or powdering occurs. The same size tower packing should be used for replacement.
The standard size of the ceramic saddle or a stainless steel pall ring is one inch.
A large carryover of liquid hydrocarbons into the glycol system can be very troublesome and dangerous. The hydrocarbons will flash in the reboiler, flood the stripper, and increase glycol losses. Heavy hydrocarbon vapors and/or liquids could also spill over the reboiler and create a serious fire hazard. Therefore, the vapors leaving the stripper vent should be piped away from the process equipment as a safety measure. The vent line should be properly sloped all the way from the stripper to the point of discharge to prevent condensed liquids from plugging the line.
Still Emissions
Vapor from the still column can contain some hydrocarbon gases that flashed from the glycol, stripping gas and aromatics. Glycol preferentially absorbs aromatics and napthene components over paraffinic components in the inlet gas.
Aromatics:
Include benzine, ethylene, toluene, and xylene (commonly called BETX)
Condense with water vapor
Could lead to “soluble” oil in the produced water discharge
Treatment consists of condensing the water vapor and BETX exiting the still column and then compressing the non-condensables (hydrocarbon gases) (Figure 5-46)
Fig. 5-46. Process flow diagram of treatment of still emissions.
5.15.6 Reboiler
The reboiler supplies heat to separate the glycol and water by simple distillation. Large plant locations may use hot oil or steam in the reboiler. Remote field locations are generally equipped with a direct-fired heaters, with the following characteristics:
Use a portion of the gas for fuel
Heating element usually has a U-Tube shape and contains one or more burners
Conservatively designed to insure long tube life and prevent glycol decomposition caused by overheating (Forming a sludge covering parts of flame “u-tube” due to glycol decomposition).
Reboiler should be equipped with a high-temperature safety overriding controller to shut down the fuel supply gas system in case of malfunction of the primary temperature controller.
The firebox heat flux (a measure of the heat transfer rate in Btu/hr/ft2) should be high enough to provide adequate heating capacity but low enough to prevent glycol decomposition. Excessive heat flux, a result of too much heat in a small area, will thermally decompose the glycol.
Flame should be correctly adjusted to give a long, rolling, and slightly yellow-tipped flame.
Nozzles are available that distribute the flame more evenly along the tube:
Decreases the heat flux of the area nearest the nozzle without actually lowering the total heat energy transferred
Avoids direct and hard impingement of the flame against the firetube
A continuous spark ignition system, or a spark igniter to relight the pilot if it goes out, is also useful.
Orifices on the air-gas mixers and pilots should be cleaned regularly to prevent burner failures.
The following temperatures in the reboiler should not be exceeded:
Type of Glycol Thermal Decomposition Temperatures
Ethylene 3290F (1650C)
Diethylene 3280F (1640C)
Triethylene 4040F (2070C)
Table.5-6. Thermal decomposition of Glycols.
Excessive discoloration and very slow degradation will result when the reboiler bulk temperature is maintained about 100F (50C) in excess of the above listed temperatures. If coke, tarry products, and/or salt deposit on the firetube, the heat transfer rate is reduced and a tube failure can result.
Localized overheating, especially where salt accumulates, will decompose the glycol.
An analysis of the glycol determines the amounts and types of these contaminants.
Salt deposits can also be detected by shutting off the burner on the reboiler at night and looking down the firebox. A bright red-glowing light will be visible at spots on the tubes where salt deposits have collected. These deposits can cause a rapid firetube burnout, particularly if the plant inlet separator is inadequate and a slug of salt water enters the absorber.
Coke and tarry products present in the circulating glycol can be removed by good filtration.
More elaborate equipment is needed to remove the salt. Contaminates, which have already deposited on the firetube and other equipment, can only be removed by using chemicals.
The heating process must be thermostatically controlled and fully automatic.
The reboiler temperature should be occasionally verified with a test thermometer to make sure true readings are being recorded.
If water and/or hydrocarbons enter the reboiler from the absorber, it may be impossible to raise the temperature until this problem is corrected. Standard orifices furnished for reboiler burners are sized for 1000–1100 Btu/scf of gas. If the rating of the fuel gas is less than this, it may be necessary to install a larger orifice or drill out the existing orifice to the next higher size.
During a unit startup, it is imperative the reboiler temperature be up to the desired operating level before flowing gas through the absorber.
The reboiler must be horizontal when erected. A nonhorizontal position can cause a firetube burnout. The reboiler should also be located close enough to the absorber to prevent excessive cooling of the lean glycol during cold weather. This will prevent hydrocarbon condensation and high glycol losses in the absorber.
5.15.7 Stripping Gas
Stripping gas is an optional item used to achieve very high glycol concentrations which cannot be obtained with normal regeneration. It will provide the maximum dew point depression and greater dehydration. Stripping gas is used to remove the residual water after the glycol has been reconcentrated in the regeneration equipment. It is used to provide intimate contact between the hot gas and the lean glycol after most of the water has been removed by distillation.
Lean glycol concentrations in the range of 99.5 - 99.9% and dew-point depressions of 1400F and above have been reported. There are several methods of introducing stripping gas into the system.
One method is to use a vertical tray or packed section in the downcomer between the reboiler and storage tank where the dry gas strips the additional water out of the regenerated glycol.
The glycol from the reboiler flows down through this section, contacts the stripping gas to remove the excess water, and goes into the storage tank.
Another method is to use glycol stripping gas sparger in the reboiler beneath the firetube.
As the glycol flows through the reboiler, gas is injected into this vessel and is heated by the glycol. Stripping gas contacts the glycol in the reboiler and removes some of the additional water.
Gas then passes out the stripper to the waste pit. The lean glycol flows from the reboiler down into the storage tank.
Stripper gas is normally taken from the reboiler fuel gas line (if dehydrated gas) at the fuel drip pot pressure. Air or oxygen should not be used. Stripping gas is usually controlled by a manual valve with a pressure gauge to indicate the flow rate through an orifice.
Stripping gas rate has the following characteristics:
Will vary according to the lean concentration desired and the method of glycol-gas contact
Usually between 2 and 10 scf/gallon of glycol circulated
Should not get high enough to flood the stripper and blow glycol out to the pit.
When stripping gas is used it is necessary to provide a more reflux in the still column to prevent excessive glycol losses. This is usually provided by using a cool glycol condenser coil in the stripper.
5.15.8 Glycol Circulating Pump
A circulating pump is used to move glycol through the system. It can be powered by electricity, gas, steam, or gas and glycol, depending upon the operating conditions and unit location.
Glycol-Gas Powered Pump
Powered by gas entrained in the wet glycol leaving the contactor. Where it utilizes the rich glycol under pressure in the absorber to furnish part of its required driving energy.
Does not require contactor glycol liquid level control, dump valve, or external power (electricity).
Gas consumption is relatively low (at 1000 psi operating pressure on the absorber, the volume of gas required is approximately 5.5 scf per gallon of lean glycol circulated)
Have few moving parts, which translates into less wear and simplified repairs.
Contact with hydrocarbon distillate, which may be entrained in glycol passing through the pump, swells o-ring seals in the pumps causing premature pump failure.
Generally used on small isolated systems
Temperatures above 200 0F damage o-ring seals.
Controls are serviceable, dependable, and, if adjusted properly, should give a long, trouble free operation, and inexpensive.
Electric Driven Positive Displacement Piston/Plunger Pump
Usually used in large installations
Require a small glycol leak in the piston rod packing for lubrication.
Resilient to hydrocarbon distillate, grit, and debris that would damage the glycol-gas powered pumps.
Pump rate should be commensurate with the gas volume being processed. Proportioning adjustments allow increased gas-glycol contact time in the absorber.
When the pump check valves become worn or clogged, the pump will operate normally except no fluid will go to the absorber. Even a pressure gauge will indicate a pumping cycle. The only evidence of this type of failure is little or no dew-point depression. One sure way to check the volume flowing is to close the valve on the absorber outlet and calculate the flow by measuring the rise in the gauge glass (if one is available) versus the amount pumped normally
One of the most common sources of glycol loss occurs at the pump packing gland. If the pump leaks over one quarts (0.25 gallon) of glycol per day, the packing needs to be replaced. An adjustment will not recover the seal. Packing should be installed hand-tight and then backed off one complete turn. If the packing gets too tight, the pistons can score and require replacement.
Glycol circulation rate of 2 to 3 gallons/lb of water to be removed is sufficient to provide adequate dehydration. An excessive rate can overload the reboiler and reduce the dehydration efficiency.
The rate should be checked regularly by timing the pump to make sure it is running at the proper speed.
Proper pump maintenance will reduce the operating costs. When the pump is not working the whole system must be shut down because the gas cannot be dried effectively without a good continuous flow of glycol in the absorber. Pumps should be lubricated regularly.
If there is insufficient glycol circulation:
Check the pump suction strainer for plugging and/or open the bleeder valve to eliminate air lock.
Glycol strainers should be regularly cleaned to avoid pump wear and other problems.
The maximum operating temperature of the pump is limited by the moving O-ring seals and nylon D slides. A maximum temperature of 2000F (940C) is recommended.
Packing life will be extended considerably if the temperature is held to a maximum of 1500F (660C). Therefore, sufficient heat exchange is necessary to keep the dry, lean glycol below these temperatures when it goes through the pump.
The pump is usually the most overworked and overused piece of equipment in the glycol process system.
The glycol system usually contains a second spare pump to avoid shutdowns when the primary pump fails. It is not uncommon for operators to use the second pump to send more glycol to the absorber to avoid wet sales gas problems. This procedure just increases operating problems. All of the other process variables should first be checked before a second pump is used.
A pressure gauge is furnished on the discharge side of the pump. Pressure gauge can be used to see that the pump is working by watching the gauge “kick” as the pump piston strokes. The sensing element in the pressure gauge is a bourdon tube. The flexing or movement of this tube indicates the pressure. A bourdon tube will fatigue or fail if subjected to continuous fluctuations in pressure on the pump discharge. Pressure should be kept off the gauge except when testing the unit or to determine glycol loss from the gauge failure.
5.15.9 Flash Tank or Glycol-Gas Separator
The flash tank, or glycol-gas separator, is an optional piece of equipment used to recover the off-gas from the glycol-powered pump and the gaseous hydrocarbons from the rich glycol. The recovered gas can be used as fuel to the reboiler and/or stripping gas. Any excess gas is usually discharged through a back pressure valve. The flash tank will keep volatile hydrocarbons out of the reboiler. The separator usually works best in a temperature range of 1300F to 1700F (550C to 770C). A two-phase separator, with at least a five minute retention time, can be used to remove the gas. If liquid hydrocarbons are present in the rich glycol, a three-phase separator should be used to remove these liquids before they get in the stripper and reboiler. A liquid retention time of 20 to 45 minutes, depending on the type of hydrocarbons, API gravity, and the amount of foam, should be provided in the vessel. Vessels should be located ahead of or behind the preheat coil in the storage tank, depending on the type of hydrocarbons present.
5.15.10 Gas Blanket
A gas blanket prevents air from contacting glycol in the reboiler and storage tanks. A small amount of low-pressure gas is bled into the storage tank. Gas is piped from the storage tank to the bottom of the stripper and it passes on overhead with the water vapor. Elimination of air helps prevent glycol decomposition by slow oxidation. The gas blanket equalizes the pressure between the reboiler and storage tank. The gas blanket also prevents the liquid seal from breaking down between these two vessels.
5.15.11 Reclaimer
The reclaimer purifies the glycol for further use by vacuum distillation. Clean glycol is driven off and all the dirty sludge is left in the vessel and then washed to the sewer. It is normally used only in very large glycol systems.
5.16 Improving Glycol Filtration
Filters will extend the life of pumps, will prevent an accumulation of solids in the absorberand in the regeneration equipment
Solids that settle out on metal surfaces will frequently set up cell corrosion. Filters remove the solids to also eliminate fouling, foaming, and plugging. Filters should be designed to remove all solid particles 5 microns and larger. They should be able to operate up to a differential pressure of 20–25 psi without loss of seal or channeling of flow. Normal differential pressure is 3-6 psi.
An internal relief valve with a setting of about 25 psi and differential pressure gauges are very helpful. New elements should be installed before the relief valve opens. When this differential increases to 15-20 psi, the element should be replaced. Filters are not usually placed in the rich glycol line, but the lean glycol can also be filtered to help keep the glycol clean. Frequent filter changes may be needed during unit startup or when neutralizers are added, to control the glycol pH.
It is important to know when and how to change elements to keep air out the glycol system. Valves and gauges should be inspected occasionally for corrosion and scale buildup. To determine the proper use of filter elements, cut them to the core and inspect them. If they are dirty throughout, the filter is being used properly. If the element is clean on the inside, an element with a different micron size may be needed. It is also a good practice to occasionally scrape some sludge from a dirty element and have it analyzed. This will help establish the types of contaminants present. A record of the number of elements replaced will establish the amount of contaminants present.
Use of Carbon Purification
Activated carbon can effectively eliminate most foaming problems by removing the hydrocarbons, well-treating chemicals, compressor oils, and other troublesome impurities from the glycol.
Two ways glycol purification can be achieved are as follows:
a- One method is to use two carbon towers installed in series but piped so they can be taken off-stream or interchanged without difficulty. In large systems about 2% of the total glycol flow should pass through the carbon towers. In small systems 100% of the total glycol flow should pass through the carbon towers. Each carbon bed should be sized to handle 2 gallons of glycol per square foot of cross-sectional area per minute. Towers should be designed to permit back flushing with water to remove the dust after the carbon is loaded. To achieve this, a retainer screen, with a smaller mesh size than the carbon should be installed above the carbon bed between the liquid inlet distributor and the outlet water drain nozzle to hold the carbon to the vessel. The liquid distributor is needed to avoid glycol channeling through the carbon. The inlet water nozzle for back-flushing should be placed below the screen in the bottom of the tower. The appearance of the glycol can generally be used to determine when the carbon needs to be regenerated or replaced. The pressure drop across the carbon bed can also be used. The pressure drop normally across the carbon bed is only 1 or 2 psi. When the pressure drop reaches 10 to 15 psi, the carbon is usually completely plugged with impurities. Steam cleaning can sometimes be used to regenerate the carbon by removing the impurities. However, this can be hazardous and offers only limited success.
b- Another method of purification is to use activated carbon in elements, such as Peco-Char.
Either purification system should be placed downstream from the solids filter. This will increase the carbon adsorptive efficiency and life.
Adsorption
5.17 Overview of Adsorption Processes
The two types of adsorption are physical adsorption and chemisorption. In physical adsorption, the bonding between the adsorbed species and the solid phase is called van der Waals forces, the attractive and repulsive forces that hold liquids and solids together and give them their structure. In chemisorption, a much stronger chemical bonding occurs between the surface and the adsorbed molecules. This chapter considers only physical adsorption, and all references to adsorption mean physical adsorption.
Adsorption is a physical phenomenon that occurs when molecules of a gas are brought into contact with a solid surface and some of them condense on the surface.
Physical adsorption is an equilibrium process like vapor−liquid equilibria. Thus, for a given vapor-phase concentration (partial pressure, and temperature), an equilibrium concentration exists on the adsorbent surface that is the maximum concentration of the condensed component (adsorbate) on the surface.
Dehydration of a gas (or a liquid hydrocarbon) with a dry desiccant is an adsorption process in which water molecules are preferentially held by the desiccant and removed from the gas stream.
Adsorption involves a form of adhesion between the surface of the solid desiccant and the water vapor in the gas. Water forms a thin film that is held to the desiccant surface by forces of attraction, not by chemical reaction.
Desiccant is a solid, granulated dehydrating medium with a large effective surface area (large number of small pores) per unit weight.
Figure 5.47 shows the equilibrium conditions for water on a commercial molecular sieve. Such curves are called isotherms. The figure is based upon a water−air mixture but is applicable to natural gas systems. The important parameter is the partial pressure of water.
The achievement of equilibrium on a small surface displays the following pattern:
Some passing molecules will condense on the surface (physical as opposed to chemical absorption).
After some finite time the molecule may acquire sufficient energy to leave and be replaced by another.
After sufficient time, a state of equilibrium will be reached wherein the number of molecules leaving the surface will equal the number arriving.
Because adsorbate concentrations are usually low, generally only a few layers of molecules will build up on the surface. Thus, adsorption processes use solids with extremely high surface-to-volume ratios. Typical desiccants (zeolite) might have as much as 4 million square feet of surface area per pound. In the case of molecular sieves, the adsorbent consists of extremely fine zeolite particles held together by a binder. Therefore, adsorbing species travel through the macropores of the binder into the micropores of the zeolite. Adsorbents such as silica gel and alumina are formed in larger particles and require no binder.
Pore openings that lead to the inside of commercial adsorbents are of molecular size.
Molecular sieves have an extremely narrow pore distribution, whereas silica gel and alumina have wide distributions. However, a molecular sieve binder, which is usually about 20% of the weight of the total adsorbent, has large pores capable of adsorbing heavier components.
Two steps are involved in adsorbing a trace gas component. The first step is to have the component contact the surface and the second step is to have it travel through the pathways inside the adsorbent. Because this process is a two-step process and the second step is relatively slow, solid adsorbents take longer to come to equilibrium with the gas phase than in absorption processes.
In addition to concentration (i.e., partial pressure for gases), two properties of the adsorbate dictate its concentration on the absorbent surface: polarity and size. Unless the adsorbent is nonpolar, which is not the case for those used in gas plants, polar molecules, like water, will be more strongly adsorbed than weakly polar or nonpolar compounds. Thus, methane is displaced by the weakly polar acid gases that are displaced by the strongly polar water.
How size affects adsorption depends upon the pore size of the adsorbent. An adsorbate too large to fit into the pores adsorbs only on the outer surface of adsorbent, which is a trivial amount of surface area compared with the pore area.
If the pores are sufficiently large to hold different adsorbates, the less volatile, which usually correlates with size, adsorbates will displace the more volatile ones. Therefore, ethane is displaced by propane.
In commercial practice, adsorption is carried out in a vertical, fixed bed of adsorbent, with the feed gas flowing down through the bed. As noted above, the process is not instantaneous, which leads to the formation of a mass transfer zone (MTZ) in the bed. Figure 5.48 shows the three zones in an adsorbent bed:
Fig. 5.47 Water loading on UOP Adsorbent 4A-DG MOLSIV Pellets. Activation conditions for the adsorbent were 662°F (350°C) and less than 10 microns Hg.
1. The equilibrium zone, where the adsorbate on the adsorbent is in equilibrium with the adsorbate in the inlet gas phase and no additional adsorption occurs
2. The mass transfer zone (MTZ), the volume where mass transfer and adsorption take place
3. The active zone, where no adsorption has yet taken place.
In the mass transfer zone (MTZ), the concentration drops from the inlet value, yin, to the outlet value, yout, in a smooth S-shaped curve. If the mass transfer rate were infinite, the MTZ would have zero thickness. The MTZ is usually assumed to form quickly in the adsorption bed and to have a constant length as it moves through the bed, unless particle size or shape is changed.
The length of the MTZ is usually 0.5 to 6 ft (0.2 to 1.8 m), and the gas is in the zone for 0.5 to 2 seconds. To maximize bed capacity, the MTZ needs to be as small as possible because the zone nominally holds only 50% of the adsorbate held by a comparable length of adsorbent at equilibrium. Both tall, slender beds, which reduce the percentage of the bed in the MTZ, and smaller particles make more of the bed effective. However, smaller particle size, deeper beds, and increased gas velocity will increase pressure drop.
http://oilprocessing.net/data/documents/V5-48.png
Fig. 5.48 Vapor-phase concentration profile of an adsorbate in the three zones of an adsorption bed.
MTZ lengths can be obtained experimentally for various materials and systems and used in graphical correlations for design purposes.
As the flow of gas continues, the MTZs move downward through the bed and water displaces all of the previously adsorbed gas until, finally, the entire bed is saturated with water vapor. When the bed is completely saturated with water vapor, the outlet gas is just as wet as the inlet gas.
Towers must be switched from the adsorption cycle to the regeneration cycle (heating and cooling) before the desiccant bed has become completely saturated.
Thus, higher velocities increase the MTZ thickness.
MTZ is a function of the following factors:
Adsorbent, adsorbent particle size, fluid velocity, fluid properties, adsorbate concentration in the entering fluid, adsorbate concentration in the adsorbent if it is not fully reactivated, temperature, pressure, and past history of the system.
5.18 Properties of Industrial Adsorbents for Dehydration
Three types of commercial adsorbents are in common use in gas processing plants:
• Silica gel, which is made of pure SiO2
• Activated alumina, which is made of Al2O3
• Molecular sieves, which are made of alkali aluminosilicates and can be altered to affect adsorption characteristics.
Table 5.7 lists the more important properties of three adsorbents compiled primarily from commercial literature. The properties are representative and vary between manufacturers.
Silica Gel Activated Alumina Molecular Sieve 4A
Shape Spherical Spherical Pellets (extruded
cylinders) and beads
Bulk density lb/ft3 (kg/m3) 49 (785) 48 (769) 40 −45 (640 − 720)
Particle size 4 − 8 mesh
5 −2 mm 7−14 mesh, 1/8-inch,
3/16-inch, 1/4-inch
diameter (3-mm,
5-mm, 6-mm) 1/16-inch,1/8-inch,1/4-inch diameter cylinders
(1.6-mm, 3.2-mm, 6-mm)
Packed bed % voids 35 35 35
Specific heat
Btu/lb-°F (kJ/kg-K) 0.25 (1.05) 0.24 (1.00) 0.24 (1.00)
Surface area m2/g 650 − 750 325 – 360 600 – 800
Pore volume cm3/g 0.36 0.5 0.28
Regeneration
temperature, °F (°C) 375 (190) 320 to 430 (160 to 220) 400 to 600 (200 to 315)
Average pore diameter (A0) 22 NA 3,4,5,10
Minimum dew point temperature of effluent, °F (°C) −80 (−60) −100 (−75) −150 (−100)
Average minimum
moisture content of
effluent gas, ppmv 5 – 10 10 – 20
lb/MMscf ~ ppmv / 21.4 0.1
Table 5.7. Representative Properties of Commercial Silica Gels, Activated Alumina, and Molecular Sieve 4A
5.19 Solid Bed Adsorption Process
Although this discussion uses molecular sieve as the example of an adsorbent to remove water, with the exception of regeneration temperatures, the basic process is the same for all gas adsorption processes. Figure 5.49 shows a schematic of a two-bed adsorber system. One bed, adsorber #1 dries gas while the other bed, adsorber #2, goes through a regeneration cycle. The wet feed goes through an inlet separator that will catch any entrained liquids before the gas enters the top of the active bed. Flow is top-down to avoid bed fluidization. The dried gas then goes through a dust filter that will catch fines before the gas exits the unit. This filter must be kept working properly, especially if the gas goes on to a cryogenic section with plate-fin heat exchangers, as dust can collect in the exchangers and reduce heat transfer and dramatically increase pressure drop.
5.20 Principles of Operation
5.20.1 Introduction
The adsorption process is a batch process, with multiple desiccant beds used in cyclic operation to dry the gas on a continuous basis. The number and arrangement of the desiccant beds may vary from two towers, adsorbing alternatively (Figure 5-49), to many towers.
Three separate functions or cycles must alternatively be performed in each dehydrator tower:
Adsorbing or gas-drying cycle
Heating or regeneration cycle
Cooling cycle (prepares the regenerated bed for another adsorbing or gas-drying cycle)
Fig. 5.49 Schematic of a two-bed adsorption unit. Valves are set to have absorber #1 in drying cycle and absorber #2 in regeneration cycle.
5.20.2 Drying Cycle
Several automatically operated switching valves and a controller route the inlet gas and
regeneration gas to the right tower at the proper time.
As the wet gas flows downward through the tower on the adsorption cycle, each of the adsorbable components is adsorbed at a different rate.
The water vapor is immediately adsorbed in the top layers of the desiccant bed.
Some of the light hydrocarbon gases and heavier hydrocarbons moving down through the bed are also adsorbed.
Heavier hydrocarbons will displace the lighter ones in the desiccant bed as the adsorbing cycle proceeds.
As the upper layers of desiccant become saturated with water, water in the wet gas stream begins displacing the previously adsorbed hydrocarbons in the lower layers.
For each component in the inlet gas stream, there will be a section of bed depth, from top to bottom, where the desiccant is saturated with that component and where the desiccant below is just starting to adsorb it.
5.20.4 Regeneration Cycle
One regeneration-gas supply scheme consists of taking a portion (5 to 15%) of the entering wet gas stream across a pressure-reducing valve that forces a portion of the upstream gas through the regeneration system. (Sales gas is sometimes used instead of a slip stream. The sales gas stream has the advantage of being free of heavier hydrocarbons that can cause coking.) The regeneration gas is heated to about 600°F (315°C) to both heat the bed and remove adsorbed water from the adsorbent. If COS formation is a problem, it can be mitigated by lowering regeneration temperatures to 400 °F (200°C) or lower, provided sufficient time for regeneration is available, or by switching to 3A. Regeneration gas enters at the bottom of the bed countercurrent to flow during adsorption to ensure that the lower part of the bed is the driest and that any contaminants trapped in the upper section of the bed stay out of the lower section.
Initially, the hot regeneration gas must heat up the tower and the desiccant.
The water begins vaporizing when the effluent hot gas temperature reaches between 240 0F and 250 0F. The bed continues to heat up slowly as the water is being desorbed or driven out of the desiccant. After all the water has been removed, heating is maintained to drive off any heavier hydrocarbons and contaminants that would not vaporize at lower temperatures.
The desiccant bed will be properly regenerated when the outlet gas (peak-out) temperature has reached between 350 0F and 550 0F.
After the heating cycle, the desiccant bed is cooled by flowing unheated regeneration gas until the desiccant is sufficiently cooled. Gas flow during this step can be concurrent or countercurrent.
All of the regeneration gas used in the heating and cooling cycles is passed through a heat exchanger (normally an aerial cooler) where it is cooled to condense the water removed from the regenerated desiccant bed.
This water is separated in the regeneration gas separator, and the gas is mixed with the incoming wet gas stream. This entire procedure is continuous and automatic.
The regeneration gas velocity is important, especially when effluent moisture contents below 1 ppm are needed (lb/MMscf ~ ppmv / 21.4). Figure 5.58 or for conservative operation velocities should not be less than 10 ft./sec., to prevent hot gases from channeling through the bed, leaving excess water in the bed after regeneration which results in poor dehydration.
Table 5.8 lists design parameters that are guidelines for typical molecular sieve dehydrators
Feed rate 10 to 1500 MMscfd (0.3 to 42 MMSm3/d)
Superficial velocity Approximately 30 to 35 ft/min (9 to 11 m/min)
Pressure drop Approximately 5 psi (35 kPa), not to exceed 10 psi (69 kPa)
Cycle time Four to 24 hours; 8 or a multiple thereof is common
Temperatures and pressures of adsorption
Temperatures and pressures of regeneration Temperatures: 50 to 115°F (10 to 45°C) Pressures: to 1500 psig (100 barg),
Temperatures: 400 to 600°F (200 to 315°C) Pressures: Adsorption pressure or lower.
Table. 5.8. Typical Operating Conditions for Molecular Sieve Dehydration Units
5.21 Adsorption System Performance
Advantages Disadvantages
- Can achieve very low dew points (-150 0F “less than 1 ppm”)
- High contact temperatures are possible
- Adaptable to large rate and load changes - High initial cost
- Batch process
- Experiences high-pressure drop through the bed
- Desiccant is sensitive to poisoning with liquids or other impurities in the gas
Table.5.9. Adsorption system performance.
Fundamentals of Oil and Gas Processing Book
Basics of Gas Field Processing Book
Prediction and Inhibition of Gas Hydrates Book
Basics of Corrosion in Oil and Gas Industry Book
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5.14 Glycol System Cleaning
Chemicals are frequently needed to clean the glycol system. If chemical cleaning is done properly, it can be quite beneficial to plant operations. While if done poorly, it can be quite costly and create long-lasting problems.
The most effective type of cleaner is a very heavy-duty alkaline solution. To provide optimum cleaning, the concentration, temperature, and pumping rate of the solution must be carefully controlled and an experienced, reputable vendor employed. A cascading technique can be used to save on the cost of cleaning chemicals.
Cleaning Techniques to Avoid
Steam cleaning is not effective and can be damaging and dangerous. It tends to harden the deposits in the system, making them almost impossible to remove.
The use of cold or hot water, with or without high detergent soaps, will do little good in cleaning the system. High-detergent soaps can create a serious problem by leaving trace quantities of soap after the cleaning job. Soap traces left in the system can make glycol foam for a long time.
Acid cleaning is good for removing inorganic deposits. Since most deposits in the glycol system are organic, acid cleaning is not very effective. It can easily create additional problems in the glycol system after the cleaning job.
5.15 Eliminating Operating Problems
Most operating problems are caused by mechanical failure. It is important to keep equipment in good working order. Following operating and maintenance suggestions helps provide a trouble-free operation.
5.15.1 Inlet Scrubber/Microfiber Filter Separator
The cleaner the inlet gas entering the absorber, the fewer operating problems there will be.
Potential problems if there was not an inlet scrubber of filter separator are:
Liquid water carryover, results in one or more of the following:
Dilutes the glycol
Lowers the absorber efficiency
Requires a greater glycol circulation rate
Increases the vapor-liquid load on the still column
Floods the still column
Vastly increases the reboiler heat load and fuel gas requirements
These problems also cause:
High glycol losses
Wet sales gases
If the water contained salt and solids, they would be deposited in the reboiler to foul the heating surfaces and possibly cause them to burn out.
The scrubber also prevents salt or other solids from entering the glycol system, where they could be deposited in the reconcentrator to:
Foul the heating surfaces
Burn out as hot spots
If liquid hydrocarbons were present:
They would pass onto the still column and reboiler
Lighter fractions would pass overhead as vapor and could create a fire hazard
Heavy fractions would collect on the glycol surface in the storage tank and could overflow the system.
Flashing of the hydrocarbon vapor can flood the still column and vastly increase the reboiler heat load and result in glycol losses.
Well corrosion control program should be planned and coordinated to prevent glycol contamination.
Scrubber or filter separator may be an integral part of the absorber or preferably a separate vessel. Vessel should be large enough to remove all solids and free liquids to keep these impurities from getting into the glycol system. Vessel should be regularly inspected to prevent any malfunction. Liquid dump line should be protected from freezing during cold weather.
Separator should be located close to the absorber so the gas does not condense more liquids before it enters the absorber.
Sometimes an efficient mist extractor, which removes all contaminants over one micron is needed between the inlet separator and the glycol plant to clean the incoming gas; this is particularly useful when paraffin and other impurities are present in a fine vapor form
When gas is compressed prior to dehydration, a coalescing type of scrubber (microfiber filter separator) placed ahead of the absorber insures removal of compressor oil in vapor form. Compressor oil and distillate can coat the tower packing either in the absorber or still column and decrease its effectiveness.
5.15.2 Absorber
This vessel contains valve or bubble cap trays or packing to give good gas-liquid contact.
Cleanliness is very important to prevent high sales gas dew points caused by foaming and/or poor gas-liquid contact. Plugged trays or packing could also increase glycol losses.
Unit startup considerations are as follows:
The pressure on the absorber should be slowly brought up to the operating range and then the glycol should be circulated to get a liquid level on all trays.
Next, the gas rate going to the absorber should be slowly increased until the operating level is reached.
If the gas enters the absorber before the trays are sealed with liquid, it will pass through the downcomers and bubble caps. When this condition occurs and the glycol is pumped into the absorber, the liquid has difficulty in sealing the downcomers, and the liquid will be carried out with the gas stream instead of flowing to the bottom of the absorber.
Gas flow rate should be increased slowly when changing from a low to a high flow rate.
Rapid surges of gas through the absorber may cause; Sufficient pressure drop through the trays to break the liquid seals, and/or Glycol to be lifted off the trays, which will flood the mist extractor and increase glycol losses
Unit shutdown considerations are as follows:
First, the fuel to the reboiler should be shut down.
Then the circulating pump should be run until the reboiler temperature is lowered to approximately 2000F (94 0C).
This precaution will prevent glycol decomposition caused by overheating.
The unit can then be shut down by slowly reducing the gas flow to prevent any unnecessary shocks on the absorber and piping.
The unit should be depressurized slowly to prevent a loss of glycol.
The dehydrator should always be depressurized from the downstream (gas outlet) side of the absorber.
A dehydrator installed on the discharge side of a compressor should be equipped with a check valve in the inlet line, located as close as possible to the absorber. Experience has shown that some glycol is sucked back into this line when a compressor backfires or is shut down.
Internal absorber damage to the trays and mesh pad may also occur with a compressor failure.
The installation of the check valve usually eliminates this problem.
All compressors taking gas from or feeding gas to a dehydrator should have pulsation dampeners. The absence of this safety device may cause fatigue failure of instruments, trays, coils, mesh pads and other parts of the dehydrator.
The glycol dump valve and level controller should be set for throttling action to give an even flow of glycol to the regenerator. This will prevent slugs, which could flood the stripper and cause excessive glycol losses.
The absorber must be vertical to insure the proper flow of glycol in the vessel and adequate contact of the glycol and gas. Inspection ports at the trays can be very useful when inspecting or cleaning the vessel.
If dry gas from a glycol unit is used for gas lift, care must be used in both sizing and operating the unit because of the unsteady gas rate required in this service. Control valves or backpressure valves are used to prevent a sudden overloading of the absorber which can break the downcomer seals in a tray type of vessel and cause excessive loss of glycol in the sales gas.
Absorbers sometimes need to be insulated when excessive condensation of light hydrocarbons collect on the vessel walls. This often occurs when dehydrating rich, warm gases in cold climates.
These very light hydrocarbons can cause tray flooding in the absorber and excessive glycol losses from the regenerator.
The mist extractor should receive special attention because glycol entrainment and well-crawling are difficult to effectively control.
The type and thickness of the mesh pad should be carefully studied to minimize glycol losses.
Care should also be taken after installation to avoid mesh pad damage. The maximum pressure drop through the contractor to avoid damage to the mesh pad is approximately 15 psi.
5.15.3 Glycol-Gas Heat Exchanger
Most units are supplied with a glycol-gas heat exchanger that uses the gas leaving the absorber to cool the lean glycol entering the absorber. This exchanger may be a coil in the top of the absorber or an external one. A water-cooled exchanger may be used when heating of the gas must be avoided. This exchanger may accumulate deposits, such as salt, solids, coke or gum which foul the heat exchanger surface, reduce the heat transfer rate, and increase the lean glycol temperature. All of the above increase glycol losses and make dehydration difficult. The vessel should be inspected regularly and cleaned when needed.
5.15.4 Lean Glycol Storage Tank or Accumulator
Normally this vessel contains a glycol heat exchanger coil which cools the lean glycol coming from the reboiler and preheats the rich glycol going to the stripper
The lean glycol is also cooled by radiation from the shell of the storage tank.
This accumulator should normally be insulated. Water cooling can also be used to help control the lean glycol temperature.
On conventional regenerators without stripping gas:
Accumulator must be vented to prevent trapping gas
Vapors, trapped in the storage tank, could cause the pump to vapor lock
A connection is usually provided in the top of the storage tank for venting
Vent line should be piped away from the process equipment and should not be connected to the stripper vent because this could cause steam to dilute the concentrated glycol
Some units are equipped to provide a dry gas blanket (no oxygen or air) in the storage tank. Blanket gas is normally piped to the regular vent connection on top of the storage tank. If blanket gas is used, it is commonly taken from the fuel gas line. Only a very slight flow of gas is required to prevent steam generated in the reboiler from contaminating the regenerated glycol.
The vessel should be inspected occasionally to see that sludge deposits and heavy hydrocarbons are not collecting in the bottom of the vessel. The heat exchanger coil should be kept clean so proper heat transfer can be made. This also prevents corrosion. If the heat exchanger develops a leak, the water rich glycol could dilute the lean glycol.
Glycol level in the storage tank should be checked and a level in the gauge glass should always be maintained. Glycol should be added as the level is pumped down. Records of the amount of glycol added should be maintained. Make certain the storage tank is not overfilled.
5.15.5 Stripper or Still Column
The stripper, or still column, is generally a packed column located on top of the reboiler to separate the water and glycol by fractional distillation. Packing is usually a ceramic saddle but 304 stainless steel pall rings can be used. A standard stripper usually has a finned atmospheric condenser in the top to cool the steam vapors and recover the entrained glycol.
Atmospheric condenser depends upon air circulation to cool the hot vapors. Increased glycol losses can occur on extremely hot days when insufficient cooling in the condenser causes poor condensation. High glycol losses can also occur on extremely cold, windy days when excessive condensation (water and glycol) overloads the reboiler. Excess liquids percolate out the stripper vent.
If stripping gas is used, an internal reflux coil is normally provided to cool the vapors to prevent excessive glycol losses. This is due to a larger mass of vapor leaving the stripper which will carry glycol. Adequate reflux is provided by passing the cool, rich glycol from the absorber through the condenser coil in the stripper. If properly adjusted, it can provide uniform condensation throughout the year.
Manual/automatic valve in the piping is furnished to bypass the reflux coil.
Under normal circumstances this valve will be closed and the total flow will be through the reflux coil. In cold weather operation, with extreme low ambient temperatures, this could produce too much reflux and the regenerator could become overloaded. Therefore, a portion or all of the rich glycol solution should bypass the reflux coil. This is accomplished by opening the manual/automatic valve until the reboiler can hold the temperature. This lowers the amount of reflux produced by the coil and reduces the load on the reboiler.
Sometimes a leak can develop in the cool glycol reflux coil in the top of the stripper. When this happens, excess glycol can flood the tower packing in the still column, upset the distillation operation, and increase glycol losses.
Broken, powdered packing can cause solution foaming in the stripper and increase glycol losses.
Packing is usually broken by excessive bed movement which is caused when hydrocarbons flash in the reboiler. Careless handling when installing the packing can also cause powdering.
As particles break down, the pressure drop through the stripper increases. This restricts the flow of vapor and liquid and causes the glycol to percolate out the top of the stripper.
Dirty packing, caused by sludge deposits of salt or tarry hydrocarbons, will also cause solution foaming in the stripper and increase glycol losses. Packing should be cleaned or replaced when plugging or powdering occurs. The same size tower packing should be used for replacement.
The standard size of the ceramic saddle or a stainless steel pall ring is one inch.
A large carryover of liquid hydrocarbons into the glycol system can be very troublesome and dangerous. The hydrocarbons will flash in the reboiler, flood the stripper, and increase glycol losses. Heavy hydrocarbon vapors and/or liquids could also spill over the reboiler and create a serious fire hazard. Therefore, the vapors leaving the stripper vent should be piped away from the process equipment as a safety measure. The vent line should be properly sloped all the way from the stripper to the point of discharge to prevent condensed liquids from plugging the line.
Still Emissions
Vapor from the still column can contain some hydrocarbon gases that flashed from the glycol, stripping gas and aromatics. Glycol preferentially absorbs aromatics and napthene components over paraffinic components in the inlet gas.
Aromatics:
Include benzine, ethylene, toluene, and xylene (commonly called BETX)
Condense with water vapor
Could lead to “soluble” oil in the produced water discharge
Treatment consists of condensing the water vapor and BETX exiting the still column and then compressing the non-condensables (hydrocarbon gases) (Figure 5-46)
Fig. 5-46. Process flow diagram of treatment of still emissions.
5.15.6 Reboiler
The reboiler supplies heat to separate the glycol and water by simple distillation. Large plant locations may use hot oil or steam in the reboiler. Remote field locations are generally equipped with a direct-fired heaters, with the following characteristics:
Use a portion of the gas for fuel
Heating element usually has a U-Tube shape and contains one or more burners
Conservatively designed to insure long tube life and prevent glycol decomposition caused by overheating (Forming a sludge covering parts of flame “u-tube” due to glycol decomposition).
Reboiler should be equipped with a high-temperature safety overriding controller to shut down the fuel supply gas system in case of malfunction of the primary temperature controller.
The firebox heat flux (a measure of the heat transfer rate in Btu/hr/ft2) should be high enough to provide adequate heating capacity but low enough to prevent glycol decomposition. Excessive heat flux, a result of too much heat in a small area, will thermally decompose the glycol.
Flame should be correctly adjusted to give a long, rolling, and slightly yellow-tipped flame.
Nozzles are available that distribute the flame more evenly along the tube:
Decreases the heat flux of the area nearest the nozzle without actually lowering the total heat energy transferred
Avoids direct and hard impingement of the flame against the firetube
A continuous spark ignition system, or a spark igniter to relight the pilot if it goes out, is also useful.
Orifices on the air-gas mixers and pilots should be cleaned regularly to prevent burner failures.
The following temperatures in the reboiler should not be exceeded:
Type of Glycol Thermal Decomposition Temperatures
Ethylene 3290F (1650C)
Diethylene 3280F (1640C)
Triethylene 4040F (2070C)
Table.5-6. Thermal decomposition of Glycols.
Excessive discoloration and very slow degradation will result when the reboiler bulk temperature is maintained about 100F (50C) in excess of the above listed temperatures. If coke, tarry products, and/or salt deposit on the firetube, the heat transfer rate is reduced and a tube failure can result.
Localized overheating, especially where salt accumulates, will decompose the glycol.
An analysis of the glycol determines the amounts and types of these contaminants.
Salt deposits can also be detected by shutting off the burner on the reboiler at night and looking down the firebox. A bright red-glowing light will be visible at spots on the tubes where salt deposits have collected. These deposits can cause a rapid firetube burnout, particularly if the plant inlet separator is inadequate and a slug of salt water enters the absorber.
Coke and tarry products present in the circulating glycol can be removed by good filtration.
More elaborate equipment is needed to remove the salt. Contaminates, which have already deposited on the firetube and other equipment, can only be removed by using chemicals.
The heating process must be thermostatically controlled and fully automatic.
The reboiler temperature should be occasionally verified with a test thermometer to make sure true readings are being recorded.
If water and/or hydrocarbons enter the reboiler from the absorber, it may be impossible to raise the temperature until this problem is corrected. Standard orifices furnished for reboiler burners are sized for 1000–1100 Btu/scf of gas. If the rating of the fuel gas is less than this, it may be necessary to install a larger orifice or drill out the existing orifice to the next higher size.
During a unit startup, it is imperative the reboiler temperature be up to the desired operating level before flowing gas through the absorber.
The reboiler must be horizontal when erected. A nonhorizontal position can cause a firetube burnout. The reboiler should also be located close enough to the absorber to prevent excessive cooling of the lean glycol during cold weather. This will prevent hydrocarbon condensation and high glycol losses in the absorber.
5.15.7 Stripping Gas
Stripping gas is an optional item used to achieve very high glycol concentrations which cannot be obtained with normal regeneration. It will provide the maximum dew point depression and greater dehydration. Stripping gas is used to remove the residual water after the glycol has been reconcentrated in the regeneration equipment. It is used to provide intimate contact between the hot gas and the lean glycol after most of the water has been removed by distillation.
Lean glycol concentrations in the range of 99.5 - 99.9% and dew-point depressions of 1400F and above have been reported. There are several methods of introducing stripping gas into the system.
One method is to use a vertical tray or packed section in the downcomer between the reboiler and storage tank where the dry gas strips the additional water out of the regenerated glycol.
The glycol from the reboiler flows down through this section, contacts the stripping gas to remove the excess water, and goes into the storage tank.
Another method is to use glycol stripping gas sparger in the reboiler beneath the firetube.
As the glycol flows through the reboiler, gas is injected into this vessel and is heated by the glycol. Stripping gas contacts the glycol in the reboiler and removes some of the additional water.
Gas then passes out the stripper to the waste pit. The lean glycol flows from the reboiler down into the storage tank.
Stripper gas is normally taken from the reboiler fuel gas line (if dehydrated gas) at the fuel drip pot pressure. Air or oxygen should not be used. Stripping gas is usually controlled by a manual valve with a pressure gauge to indicate the flow rate through an orifice.
Stripping gas rate has the following characteristics:
Will vary according to the lean concentration desired and the method of glycol-gas contact
Usually between 2 and 10 scf/gallon of glycol circulated
Should not get high enough to flood the stripper and blow glycol out to the pit.
When stripping gas is used it is necessary to provide a more reflux in the still column to prevent excessive glycol losses. This is usually provided by using a cool glycol condenser coil in the stripper.
5.15.8 Glycol Circulating Pump
A circulating pump is used to move glycol through the system. It can be powered by electricity, gas, steam, or gas and glycol, depending upon the operating conditions and unit location.
Glycol-Gas Powered Pump
Powered by gas entrained in the wet glycol leaving the contactor. Where it utilizes the rich glycol under pressure in the absorber to furnish part of its required driving energy.
Does not require contactor glycol liquid level control, dump valve, or external power (electricity).
Gas consumption is relatively low (at 1000 psi operating pressure on the absorber, the volume of gas required is approximately 5.5 scf per gallon of lean glycol circulated)
Have few moving parts, which translates into less wear and simplified repairs.
Contact with hydrocarbon distillate, which may be entrained in glycol passing through the pump, swells o-ring seals in the pumps causing premature pump failure.
Generally used on small isolated systems
Temperatures above 200 0F damage o-ring seals.
Controls are serviceable, dependable, and, if adjusted properly, should give a long, trouble free operation, and inexpensive.
Electric Driven Positive Displacement Piston/Plunger Pump
Usually used in large installations
Require a small glycol leak in the piston rod packing for lubrication.
Resilient to hydrocarbon distillate, grit, and debris that would damage the glycol-gas powered pumps.
Pump rate should be commensurate with the gas volume being processed. Proportioning adjustments allow increased gas-glycol contact time in the absorber.
When the pump check valves become worn or clogged, the pump will operate normally except no fluid will go to the absorber. Even a pressure gauge will indicate a pumping cycle. The only evidence of this type of failure is little or no dew-point depression. One sure way to check the volume flowing is to close the valve on the absorber outlet and calculate the flow by measuring the rise in the gauge glass (if one is available) versus the amount pumped normally
One of the most common sources of glycol loss occurs at the pump packing gland. If the pump leaks over one quarts (0.25 gallon) of glycol per day, the packing needs to be replaced. An adjustment will not recover the seal. Packing should be installed hand-tight and then backed off one complete turn. If the packing gets too tight, the pistons can score and require replacement.
Glycol circulation rate of 2 to 3 gallons/lb of water to be removed is sufficient to provide adequate dehydration. An excessive rate can overload the reboiler and reduce the dehydration efficiency.
The rate should be checked regularly by timing the pump to make sure it is running at the proper speed.
Proper pump maintenance will reduce the operating costs. When the pump is not working the whole system must be shut down because the gas cannot be dried effectively without a good continuous flow of glycol in the absorber. Pumps should be lubricated regularly.
If there is insufficient glycol circulation:
Check the pump suction strainer for plugging and/or open the bleeder valve to eliminate air lock.
Glycol strainers should be regularly cleaned to avoid pump wear and other problems.
The maximum operating temperature of the pump is limited by the moving O-ring seals and nylon D slides. A maximum temperature of 2000F (940C) is recommended.
Packing life will be extended considerably if the temperature is held to a maximum of 1500F (660C). Therefore, sufficient heat exchange is necessary to keep the dry, lean glycol below these temperatures when it goes through the pump.
The pump is usually the most overworked and overused piece of equipment in the glycol process system.
The glycol system usually contains a second spare pump to avoid shutdowns when the primary pump fails. It is not uncommon for operators to use the second pump to send more glycol to the absorber to avoid wet sales gas problems. This procedure just increases operating problems. All of the other process variables should first be checked before a second pump is used.
A pressure gauge is furnished on the discharge side of the pump. Pressure gauge can be used to see that the pump is working by watching the gauge “kick” as the pump piston strokes. The sensing element in the pressure gauge is a bourdon tube. The flexing or movement of this tube indicates the pressure. A bourdon tube will fatigue or fail if subjected to continuous fluctuations in pressure on the pump discharge. Pressure should be kept off the gauge except when testing the unit or to determine glycol loss from the gauge failure.
5.15.9 Flash Tank or Glycol-Gas Separator
The flash tank, or glycol-gas separator, is an optional piece of equipment used to recover the off-gas from the glycol-powered pump and the gaseous hydrocarbons from the rich glycol. The recovered gas can be used as fuel to the reboiler and/or stripping gas. Any excess gas is usually discharged through a back pressure valve. The flash tank will keep volatile hydrocarbons out of the reboiler. The separator usually works best in a temperature range of 1300F to 1700F (550C to 770C). A two-phase separator, with at least a five minute retention time, can be used to remove the gas. If liquid hydrocarbons are present in the rich glycol, a three-phase separator should be used to remove these liquids before they get in the stripper and reboiler. A liquid retention time of 20 to 45 minutes, depending on the type of hydrocarbons, API gravity, and the amount of foam, should be provided in the vessel. Vessels should be located ahead of or behind the preheat coil in the storage tank, depending on the type of hydrocarbons present.
5.15.10 Gas Blanket
A gas blanket prevents air from contacting glycol in the reboiler and storage tanks. A small amount of low-pressure gas is bled into the storage tank. Gas is piped from the storage tank to the bottom of the stripper and it passes on overhead with the water vapor. Elimination of air helps prevent glycol decomposition by slow oxidation. The gas blanket equalizes the pressure between the reboiler and storage tank. The gas blanket also prevents the liquid seal from breaking down between these two vessels.
5.15.11 Reclaimer
The reclaimer purifies the glycol for further use by vacuum distillation. Clean glycol is driven off and all the dirty sludge is left in the vessel and then washed to the sewer. It is normally used only in very large glycol systems.
5.16 Improving Glycol Filtration
Filters will extend the life of pumps, will prevent an accumulation of solids in the absorberand in the regeneration equipment
Solids that settle out on metal surfaces will frequently set up cell corrosion. Filters remove the solids to also eliminate fouling, foaming, and plugging. Filters should be designed to remove all solid particles 5 microns and larger. They should be able to operate up to a differential pressure of 20–25 psi without loss of seal or channeling of flow. Normal differential pressure is 3-6 psi.
An internal relief valve with a setting of about 25 psi and differential pressure gauges are very helpful. New elements should be installed before the relief valve opens. When this differential increases to 15-20 psi, the element should be replaced. Filters are not usually placed in the rich glycol line, but the lean glycol can also be filtered to help keep the glycol clean. Frequent filter changes may be needed during unit startup or when neutralizers are added, to control the glycol pH.
It is important to know when and how to change elements to keep air out the glycol system. Valves and gauges should be inspected occasionally for corrosion and scale buildup. To determine the proper use of filter elements, cut them to the core and inspect them. If they are dirty throughout, the filter is being used properly. If the element is clean on the inside, an element with a different micron size may be needed. It is also a good practice to occasionally scrape some sludge from a dirty element and have it analyzed. This will help establish the types of contaminants present. A record of the number of elements replaced will establish the amount of contaminants present.
Use of Carbon Purification
Activated carbon can effectively eliminate most foaming problems by removing the hydrocarbons, well-treating chemicals, compressor oils, and other troublesome impurities from the glycol.
Two ways glycol purification can be achieved are as follows:
a- One method is to use two carbon towers installed in series but piped so they can be taken off-stream or interchanged without difficulty. In large systems about 2% of the total glycol flow should pass through the carbon towers. In small systems 100% of the total glycol flow should pass through the carbon towers. Each carbon bed should be sized to handle 2 gallons of glycol per square foot of cross-sectional area per minute. Towers should be designed to permit back flushing with water to remove the dust after the carbon is loaded. To achieve this, a retainer screen, with a smaller mesh size than the carbon should be installed above the carbon bed between the liquid inlet distributor and the outlet water drain nozzle to hold the carbon to the vessel. The liquid distributor is needed to avoid glycol channeling through the carbon. The inlet water nozzle for back-flushing should be placed below the screen in the bottom of the tower. The appearance of the glycol can generally be used to determine when the carbon needs to be regenerated or replaced. The pressure drop across the carbon bed can also be used. The pressure drop normally across the carbon bed is only 1 or 2 psi. When the pressure drop reaches 10 to 15 psi, the carbon is usually completely plugged with impurities. Steam cleaning can sometimes be used to regenerate the carbon by removing the impurities. However, this can be hazardous and offers only limited success.
b- Another method of purification is to use activated carbon in elements, such as Peco-Char.
Either purification system should be placed downstream from the solids filter. This will increase the carbon adsorptive efficiency and life.
Adsorption
5.17 Overview of Adsorption Processes
The two types of adsorption are physical adsorption and chemisorption. In physical adsorption, the bonding between the adsorbed species and the solid phase is called van der Waals forces, the attractive and repulsive forces that hold liquids and solids together and give them their structure. In chemisorption, a much stronger chemical bonding occurs between the surface and the adsorbed molecules. This chapter considers only physical adsorption, and all references to adsorption mean physical adsorption.
Adsorption is a physical phenomenon that occurs when molecules of a gas are brought into contact with a solid surface and some of them condense on the surface.
Physical adsorption is an equilibrium process like vapor−liquid equilibria. Thus, for a given vapor-phase concentration (partial pressure, and temperature), an equilibrium concentration exists on the adsorbent surface that is the maximum concentration of the condensed component (adsorbate) on the surface.
Dehydration of a gas (or a liquid hydrocarbon) with a dry desiccant is an adsorption process in which water molecules are preferentially held by the desiccant and removed from the gas stream.
Adsorption involves a form of adhesion between the surface of the solid desiccant and the water vapor in the gas. Water forms a thin film that is held to the desiccant surface by forces of attraction, not by chemical reaction.
Desiccant is a solid, granulated dehydrating medium with a large effective surface area (large number of small pores) per unit weight.
Figure 5.47 shows the equilibrium conditions for water on a commercial molecular sieve. Such curves are called isotherms. The figure is based upon a water−air mixture but is applicable to natural gas systems. The important parameter is the partial pressure of water.
The achievement of equilibrium on a small surface displays the following pattern:
Some passing molecules will condense on the surface (physical as opposed to chemical absorption).
After some finite time the molecule may acquire sufficient energy to leave and be replaced by another.
After sufficient time, a state of equilibrium will be reached wherein the number of molecules leaving the surface will equal the number arriving.
Because adsorbate concentrations are usually low, generally only a few layers of molecules will build up on the surface. Thus, adsorption processes use solids with extremely high surface-to-volume ratios. Typical desiccants (zeolite) might have as much as 4 million square feet of surface area per pound. In the case of molecular sieves, the adsorbent consists of extremely fine zeolite particles held together by a binder. Therefore, adsorbing species travel through the macropores of the binder into the micropores of the zeolite. Adsorbents such as silica gel and alumina are formed in larger particles and require no binder.
Pore openings that lead to the inside of commercial adsorbents are of molecular size.
Molecular sieves have an extremely narrow pore distribution, whereas silica gel and alumina have wide distributions. However, a molecular sieve binder, which is usually about 20% of the weight of the total adsorbent, has large pores capable of adsorbing heavier components.
Two steps are involved in adsorbing a trace gas component. The first step is to have the component contact the surface and the second step is to have it travel through the pathways inside the adsorbent. Because this process is a two-step process and the second step is relatively slow, solid adsorbents take longer to come to equilibrium with the gas phase than in absorption processes.
In addition to concentration (i.e., partial pressure for gases), two properties of the adsorbate dictate its concentration on the absorbent surface: polarity and size. Unless the adsorbent is nonpolar, which is not the case for those used in gas plants, polar molecules, like water, will be more strongly adsorbed than weakly polar or nonpolar compounds. Thus, methane is displaced by the weakly polar acid gases that are displaced by the strongly polar water.
How size affects adsorption depends upon the pore size of the adsorbent. An adsorbate too large to fit into the pores adsorbs only on the outer surface of adsorbent, which is a trivial amount of surface area compared with the pore area.
If the pores are sufficiently large to hold different adsorbates, the less volatile, which usually correlates with size, adsorbates will displace the more volatile ones. Therefore, ethane is displaced by propane.
In commercial practice, adsorption is carried out in a vertical, fixed bed of adsorbent, with the feed gas flowing down through the bed. As noted above, the process is not instantaneous, which leads to the formation of a mass transfer zone (MTZ) in the bed. Figure 5.48 shows the three zones in an adsorbent bed:
Fig. 5.47 Water loading on UOP Adsorbent 4A-DG MOLSIV Pellets. Activation conditions for the adsorbent were 662°F (350°C) and less than 10 microns Hg.
1. The equilibrium zone, where the adsorbate on the adsorbent is in equilibrium with the adsorbate in the inlet gas phase and no additional adsorption occurs
2. The mass transfer zone (MTZ), the volume where mass transfer and adsorption take place
3. The active zone, where no adsorption has yet taken place.
In the mass transfer zone (MTZ), the concentration drops from the inlet value, yin, to the outlet value, yout, in a smooth S-shaped curve. If the mass transfer rate were infinite, the MTZ would have zero thickness. The MTZ is usually assumed to form quickly in the adsorption bed and to have a constant length as it moves through the bed, unless particle size or shape is changed.
The length of the MTZ is usually 0.5 to 6 ft (0.2 to 1.8 m), and the gas is in the zone for 0.5 to 2 seconds. To maximize bed capacity, the MTZ needs to be as small as possible because the zone nominally holds only 50% of the adsorbate held by a comparable length of adsorbent at equilibrium. Both tall, slender beds, which reduce the percentage of the bed in the MTZ, and smaller particles make more of the bed effective. However, smaller particle size, deeper beds, and increased gas velocity will increase pressure drop.
http://oilprocessing.net/data/documents/V5-48.png
Fig. 5.48 Vapor-phase concentration profile of an adsorbate in the three zones of an adsorption bed.
MTZ lengths can be obtained experimentally for various materials and systems and used in graphical correlations for design purposes.
As the flow of gas continues, the MTZs move downward through the bed and water displaces all of the previously adsorbed gas until, finally, the entire bed is saturated with water vapor. When the bed is completely saturated with water vapor, the outlet gas is just as wet as the inlet gas.
Towers must be switched from the adsorption cycle to the regeneration cycle (heating and cooling) before the desiccant bed has become completely saturated.
Thus, higher velocities increase the MTZ thickness.
MTZ is a function of the following factors:
Adsorbent, adsorbent particle size, fluid velocity, fluid properties, adsorbate concentration in the entering fluid, adsorbate concentration in the adsorbent if it is not fully reactivated, temperature, pressure, and past history of the system.
5.18 Properties of Industrial Adsorbents for Dehydration
Three types of commercial adsorbents are in common use in gas processing plants:
• Silica gel, which is made of pure SiO2
• Activated alumina, which is made of Al2O3
• Molecular sieves, which are made of alkali aluminosilicates and can be altered to affect adsorption characteristics.
Table 5.7 lists the more important properties of three adsorbents compiled primarily from commercial literature. The properties are representative and vary between manufacturers.
Silica Gel Activated Alumina Molecular Sieve 4A
Shape Spherical Spherical Pellets (extruded
cylinders) and beads
Bulk density lb/ft3 (kg/m3) 49 (785) 48 (769) 40 −45 (640 − 720)
Particle size 4 − 8 mesh
5 −2 mm 7−14 mesh, 1/8-inch,
3/16-inch, 1/4-inch
diameter (3-mm,
5-mm, 6-mm) 1/16-inch,1/8-inch,1/4-inch diameter cylinders
(1.6-mm, 3.2-mm, 6-mm)
Packed bed % voids 35 35 35
Specific heat
Btu/lb-°F (kJ/kg-K) 0.25 (1.05) 0.24 (1.00) 0.24 (1.00)
Surface area m2/g 650 − 750 325 – 360 600 – 800
Pore volume cm3/g 0.36 0.5 0.28
Regeneration
temperature, °F (°C) 375 (190) 320 to 430 (160 to 220) 400 to 600 (200 to 315)
Average pore diameter (A0) 22 NA 3,4,5,10
Minimum dew point temperature of effluent, °F (°C) −80 (−60) −100 (−75) −150 (−100)
Average minimum
moisture content of
effluent gas, ppmv 5 – 10 10 – 20
lb/MMscf ~ ppmv / 21.4 0.1
Table 5.7. Representative Properties of Commercial Silica Gels, Activated Alumina, and Molecular Sieve 4A
5.19 Solid Bed Adsorption Process
Although this discussion uses molecular sieve as the example of an adsorbent to remove water, with the exception of regeneration temperatures, the basic process is the same for all gas adsorption processes. Figure 5.49 shows a schematic of a two-bed adsorber system. One bed, adsorber #1 dries gas while the other bed, adsorber #2, goes through a regeneration cycle. The wet feed goes through an inlet separator that will catch any entrained liquids before the gas enters the top of the active bed. Flow is top-down to avoid bed fluidization. The dried gas then goes through a dust filter that will catch fines before the gas exits the unit. This filter must be kept working properly, especially if the gas goes on to a cryogenic section with plate-fin heat exchangers, as dust can collect in the exchangers and reduce heat transfer and dramatically increase pressure drop.
5.20 Principles of Operation
5.20.1 Introduction
The adsorption process is a batch process, with multiple desiccant beds used in cyclic operation to dry the gas on a continuous basis. The number and arrangement of the desiccant beds may vary from two towers, adsorbing alternatively (Figure 5-49), to many towers.
Three separate functions or cycles must alternatively be performed in each dehydrator tower:
Adsorbing or gas-drying cycle
Heating or regeneration cycle
Cooling cycle (prepares the regenerated bed for another adsorbing or gas-drying cycle)
Fig. 5.49 Schematic of a two-bed adsorption unit. Valves are set to have absorber #1 in drying cycle and absorber #2 in regeneration cycle.
5.20.2 Drying Cycle
Several automatically operated switching valves and a controller route the inlet gas and
regeneration gas to the right tower at the proper time.
As the wet gas flows downward through the tower on the adsorption cycle, each of the adsorbable components is adsorbed at a different rate.
The water vapor is immediately adsorbed in the top layers of the desiccant bed.
Some of the light hydrocarbon gases and heavier hydrocarbons moving down through the bed are also adsorbed.
Heavier hydrocarbons will displace the lighter ones in the desiccant bed as the adsorbing cycle proceeds.
As the upper layers of desiccant become saturated with water, water in the wet gas stream begins displacing the previously adsorbed hydrocarbons in the lower layers.
For each component in the inlet gas stream, there will be a section of bed depth, from top to bottom, where the desiccant is saturated with that component and where the desiccant below is just starting to adsorb it.
5.20.4 Regeneration Cycle
One regeneration-gas supply scheme consists of taking a portion (5 to 15%) of the entering wet gas stream across a pressure-reducing valve that forces a portion of the upstream gas through the regeneration system. (Sales gas is sometimes used instead of a slip stream. The sales gas stream has the advantage of being free of heavier hydrocarbons that can cause coking.) The regeneration gas is heated to about 600°F (315°C) to both heat the bed and remove adsorbed water from the adsorbent. If COS formation is a problem, it can be mitigated by lowering regeneration temperatures to 400 °F (200°C) or lower, provided sufficient time for regeneration is available, or by switching to 3A. Regeneration gas enters at the bottom of the bed countercurrent to flow during adsorption to ensure that the lower part of the bed is the driest and that any contaminants trapped in the upper section of the bed stay out of the lower section.
Initially, the hot regeneration gas must heat up the tower and the desiccant.
The water begins vaporizing when the effluent hot gas temperature reaches between 240 0F and 250 0F. The bed continues to heat up slowly as the water is being desorbed or driven out of the desiccant. After all the water has been removed, heating is maintained to drive off any heavier hydrocarbons and contaminants that would not vaporize at lower temperatures.
The desiccant bed will be properly regenerated when the outlet gas (peak-out) temperature has reached between 350 0F and 550 0F.
After the heating cycle, the desiccant bed is cooled by flowing unheated regeneration gas until the desiccant is sufficiently cooled. Gas flow during this step can be concurrent or countercurrent.
All of the regeneration gas used in the heating and cooling cycles is passed through a heat exchanger (normally an aerial cooler) where it is cooled to condense the water removed from the regenerated desiccant bed.
This water is separated in the regeneration gas separator, and the gas is mixed with the incoming wet gas stream. This entire procedure is continuous and automatic.
The regeneration gas velocity is important, especially when effluent moisture contents below 1 ppm are needed (lb/MMscf ~ ppmv / 21.4). Figure 5.58 or for conservative operation velocities should not be less than 10 ft./sec., to prevent hot gases from channeling through the bed, leaving excess water in the bed after regeneration which results in poor dehydration.
Table 5.8 lists design parameters that are guidelines for typical molecular sieve dehydrators
Feed rate 10 to 1500 MMscfd (0.3 to 42 MMSm3/d)
Superficial velocity Approximately 30 to 35 ft/min (9 to 11 m/min)
Pressure drop Approximately 5 psi (35 kPa), not to exceed 10 psi (69 kPa)
Cycle time Four to 24 hours; 8 or a multiple thereof is common
Temperatures and pressures of adsorption
Temperatures and pressures of regeneration Temperatures: 50 to 115°F (10 to 45°C) Pressures: to 1500 psig (100 barg),
Temperatures: 400 to 600°F (200 to 315°C) Pressures: Adsorption pressure or lower.
Table. 5.8. Typical Operating Conditions for Molecular Sieve Dehydration Units
5.21 Adsorption System Performance
Advantages Disadvantages
- Can achieve very low dew points (-150 0F “less than 1 ppm”)
- High contact temperatures are possible
- Adaptable to large rate and load changes - High initial cost
- Batch process
- Experiences high-pressure drop through the bed
- Desiccant is sensitive to poisoning with liquids or other impurities in the gas
Table.5.9. Adsorption system performance.
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Re: Basics of Gas Field Processing Book "Full text"
Chapter 5 - Part 3 B
5.22 Effect of Process Variables
For the commonly used 4A molecular sieve, the Engineering Data Book suggests that the design water content of a molecular sieve when at equilibrium with saturated gas at 75°F (24°C) will be 13 lb H2O/100 lb sieve compared with a new molecular sieve, which holds about 20 lb H2O/100 lb sieve. Two main factors affect this number: water content of entering gas and adsorption temperature.
5.22.1 Quality of Inlet Gas
Performance of dry bed dehydrator is affected by moisture content and components in the natural gas stream. The relative saturation of the inlet gas determines the size of a given desiccant bed, and affects the transfer of water to the adsorbent.
If the gas comes to the dehydration unit fully saturated, which is often the case, cooling the gas and removing the condensed water before the gas enters the bed lowers water loading.
Compounds in produced natural gas adversely affect performance of the dry bed dehydrator.
Components of concern are carbon dioxide, heavy hydrocarbons, and sulfur-bearing compounds
The greater the molecular weight of a compound, the greater its adsorption potential.
Trace amounts of oxygen affect bed life and performance in a variety of ways. At the normal regeneration gas temperature of 600°F for molecular sieves, 2 moles of oxygen react with methane to form 2 moles of water and 1 mole of CO2. As this reaction is exothermic, higher amounts of oxygen in the gas can lead to temperatures above the design temperature of the molecular sieve vessel. When oxygen is present, the temperature of the beds during regeneration must be monitored for safety reasons. Oxygen undergoes partial oxidation reactions with heavier hydrocarbons, which are adsorbed in the binder, to form alcohols and carbocylic acids that ultimately turn to water and CO2. If H2S is present, it undergoes oxidation to elemental sulfur, sulfur dioxide, and water. Oxygen concentrations greater than 20 ppmv also generate olefins that become coke in the bed. These reactions reduce molecular sieve capacity by forming solid deposits and by causing incomplete removal of water during regeneration because the partial pressure of water is higher (see Figure 5.47).
To avoid the above reactions as well as COS formation, regeneration temperatures are lowered to the 300 to 375°F range. However, this range increases the required regeneration time and the amount of regeneration gas used, which increases recompression cost.
As in all processes, ensuring that the beds are protected from entrained water and hydrocarbons is important. Even trace amounts of entrained water load the bed quickly and increase the regeneration heat load.
5.22.2 Temperature
Operation is very sensitive to the temperature of the incoming gas.
Adsorption efficiency decreases as the temperature increases.
Molecular sieves and most other adsorbents have significantly higher adsorptive capacity at low temperatures. Cooling the gas to lower temperature in order to improve adsorption is limited due to condensation and/or hydrate formation.
Temperature of the regeneration gas that commingles with the incoming wet gas ahead of the dehydrators is important. The temperature must remain within 10 0F to 15 0F, otherwise liquid water and hydrocarbons will condense as the hotter gas stream cools. Condensed liquids that strike the bed can shorten the solid desiccant’s life.
Temperature of the hot gas entering and leaving a desiccant tower during the heating cycle affects plant efficiency and the desiccant life. High regeneration gas temperature assures good removal (desorption) of water and contaminants from the bed. 450 0F to 600 0F is usually used as a regeneration temperature.
Desiccant bed temperature reached during the cooling cycle is important.
If wet gas is used to cool the desiccant: Terminate the cooling cycle when the bed reaches 125 0F. Additional cooling may cause water to be adsorbed from the wet gas stream and preload (presaturate) the bed before the next adsorption cycle begins. If dry gas is used to cool the desiccant: Terminate the cooling cycle within 10 0F to 20 0F of the incoming gas temperature. It maximizes adsorption capacity of the bed.
The temperature of the regeneration gas going through the regeneration gas scrubber should be held low enough to condense and remove the water and hydrocarbons without causing hydrate problems.
5.22.3 Pressure
The adsorption capacity of a dry bed unit decreases as pressure is lowered and with usage.
Operating dry bed dehydrators well below the design pressure requires the desiccant to work harder to remove the additional water, and maintain the desired effluent dew point. With the same volume of incoming gas, the increased gas velocity occurring at the lower pressure could affect the effluent moisture content, and damage the desiccant.
At pressure above 1300 to 1400 psia, the co-adsorption effects of hydrocarbons are very significant.
5.22.4 Cycle Time
In principle, beds can be run until the first sign of breakthrough. This practice maximizes cycle time, which extends bed life because temperature cycling is a major source of bed degeneration, and minimizes regeneration costs. However, most plants operate on a set time cycle to ensure no adsorbate breakthrough. Adsorbent capacity is not a fixed value and declines with usage. For the first few months of operation, a new desiccant normally has a high capacity for water removal. If a moisture analyzer is used on the effluent gas, a much longer drying cycle can be achieved. As the desiccant ages, the cycle time can be shortened to save regeneration fuel costs and improve the desiccant life. Common cycle times are as follows: 8 hours on stream - 5 to 6 hours heating - 2 to 3 hours cooling
5.22.5 Gas Velocities
As the gas velocity during the drying cycle decreases, the ability of the desiccant to dehydrate the gas increases. On the surface, it would seem desirable to operate at minimum flow rates to utilize the desiccant fully.
Low linear velocities:
Require towers with large cross-sectional areas to handle a given gas flow
Allow wet gas to channel through the desiccant bed and thus not be properly dehydrated.
Compromise must be made between the tower diameter and the maximum utilization of the desiccant.
High linear velocities:
Lower adsorption efficiency
May cause desiccant damage
Higher inlet compression discharge pressures to maintain the same refrigeration requirements and outlet pressure.
Increased mechanical load on the adsorbent, which leads to particle breakdown and causes further increases in pressure drop.
5.22.6 Source of Regeneration Gas
Source of regeneration gas depends on plant requirements and the availability of a suitable gas stream. Regeneration gas should be dry when low effluent moisture contents (in the range of 0.1 ppm) are required. Plant tail gate gas can normally be used.
If only moderate drying is required, a portion of the wet feed gas can be used.
Figure 5-52 is an equilibrium diagram showing lines of constant water loading. For example:
A molecular sieve bed at 100 0F in equilibrium with a gas having a -80 0F water dew point will contain about 4 wt.% water. Equilibrium curves for a given adsorbate-adsorbent can be used to estimate the regeneration conditions necessary to provide the required outlet conditions. For example, If the regeneration gas is taken from inlet gas with a dew point of 40 0F and is heated to 450 0F, the mol sieve will contain 3 wt.% water after regeneration.
If the gas to be treated is at 100 0F, the intersection of the 3 wt.% line with an adsorbent temperature of 100 0F gives the minimum attainable dew point at -95 0F.
If this dew point is not satisfactory, either the regeneration gas must be heated to above 450 0F or a gas of a higher dew point (e.g., residue gas) must be used for regeneration gas.
Fig. 5-52 Equilibrium diagram showing lines of constant water loading for a type 4a molecular sieve.
5.22.7 Direction of Gas Flow
Direction of flow during the drying cycle is downward, which:
Permits higher velocities without lifting or fluidizing the desiccant bed (fluidization can severely damage the desiccant)
Direction of flow during the heating cycle is counter-current to the direction of the adsorption flow.
It permits better reactivation of the lower portion of the desiccant bed, which must perform the super-dehydration during the drying cycle, especially in cryogenic plants.
If flow is co-current, all water and/or other contaminants must move through the entire bed, thus causing additional desiccant contamination and requiring longer regeneration times.
Direction of flow during the cooling cycle:
When dry gas is used, the flow direction is counter current to the adsorption flow, thus simplifying piping and valve configuration.
When wet gas is used, the flow direction is in the same direction as the adsorption flow so that the water adsorbed during the cooling cycle as the desiccant cools will preload on the inlet end of the bed.
If counter current flow is used in cooling with wet gas, water is deposited on the exit end of the bed. When the next adsorption cycle begins, the wet gas is immediately dried, but as the dry gas continues to move down through the bed, it picks up some of the water deposited during the cooling cycle and sometimes puts too much moisture in the effluent stream.
If wet gas is used, the additional water load, deposited during the cooling cycle, should be included when the amount of desiccant needed for dehydration is calculated.
5.22.8 Desiccant Selection
Desiccant selection is based upon:
Economics
Process conditions
Desiccants are usually interchangeable. Equipment designed for one desiccant can often operate effectively with another. No desiccant product will remain effective with massive liquid carryovers.
All desiccants are damaged by heavy impurities carried into the bed with gases. These include:
Crude oil and condensate
Glycols and amines
Most corrosion inhibitors
Well treating fluids
All desiccants exhibit a decrease in capacity (design loading) with an increase in temperature.
Molecular sieves are less affected and Aluminas are most affected.
Aluminas and molecular sieves act as a catalyst with H2S to form COS, which deposits sulfur on the desiccant bed during regeneration. (Carbonyl sulfide (COS) is formed in the following reaction: H2S + CO2 ↔ H2O + COS. Its concentration in feed gas is normally extremely low.)
Alumina gels, activated aluminas, and molecular sieves are all chemically attacked by strong mineral acids and thus decrease their adsorptive capacity.
Table 5-7 provides certain physical characteristics of the more common solid desiccants.
5.22.8.1 Molecular Sieves
Molecular sieves are a class of aluminosilicates. They produce the lowest water dewpoints, and can be used to simultaneously sweeten and dry gases and liquids. Their equilibrium water capacity is much less dependent on adsorption temperature and relative humidity. They are usually more expensive.
Molecular sieves offer the highest adsorptive capacity of all desiccants when the feed gas is at very high temperatures or at low relative saturation. It is the only desiccants capable of dehydrating gas to less than 1 ppm of water content are required for cryogenic temperatures (lb/MMscf ~ ppmv / 21.4) (dew points down to -150 0F). Therefore, for gas going into cryogenic processing, the only adsorbent that can obtain the required dehydration is a molecular sieve. Of these, 4A is the most common, but the smaller pore 3A is sometimes used. It has the advantage of being a poorer catalyst for generation of COS if both H2S and CO2 are present because a portion of the more active sodium cations in 4A has been replaced with potassium. (Carbonyl sulfide (COS) is formed in the following reaction: H2S + CO2 ↔ H2O + COS. The equilibrium constant for the reaction is of the order of magnitude of 10−6 at adsorption temperatures but increases to 10−4 at regeneration temperatures.)
If both oxygen and H2S are present 3A reduces the production of elemental sulfur that can block adsorbent pores. However, plant operators usually have little incentive to use 3A for dehydrating gas going to hydrocarbon recovery.
5.22.8.2 Silica Gel
Silica Gel and Alumina are generally offer a lower first cost.
Silica Gel is essentially pure silicon dioxide, SiO2. It is used for gas and liquid dehydration and hydrocarbon (iC5+) recovery from natural gas. When used for hydrocarbon removal, the units are often called HRUs (Hydrocarbon Recovery Units) or SCUs (Short Cycle Units). When used for dehydration, silica gel will give outlet dewpoints of approximately –60°F.
Silica gel can be regenerated to a lower water content than molecular sieves and at much lower temperatures (400 0F for gels versus 500 0 to 600 0F for sieves).
It shatters in the presence of free water or light hydrocarbon liquids. The problem is minimized by using a 4 to 6 inch buffer bed of mullite ball (or equivalent) to protect the silica gel from direct contact. Silica gels are used mostly where a high concentration of water (>1 mol%) vapor is present in the feed, and low levels of water in the dehydrated gas are not needed. They are relatively noncatalytic compounds.
5.22.8.3 Aluminas
Alumina is a hydrated form of alumina oxide (Al2O3). It is used for gas and liquid dehydration and will give outlet dewpoints of about –90°F. Less heat is required to regenerate alumina and silica gel than for molecular sieve, and the regeneration temperature is lower. Molecular sieves give lower outlet water dewpoints. Aluminas are very polar and strongly attract water and acid gases. They are used for moderate levels of water in the feed when low levels of water in the product are not required. They have the highest mechanical strength of the adsorbents considered here.
5.22.8.4 Desirable Characteristics of Solid Desiccants
High adsorptive capacity (lb/lb), which reduces contactor size.
Easy regeneration, for simplicity and economics of operation.
High rate of adsorption, which allows higher gas velocities and thereby reduces contactor size.
Low resistance to gas flow, to minimize gas pressure drop through the unit.
High adsorptive capacity retained after repeated regeneration, allowing smaller initial charge and longer service before replacement.
High mechanical strength, to resist crushing and dust formation.
Inert chemicals, to prevent chemical reactions during adsorption and regeneration.
Volume unchanged when product is wet, which would otherwise necessitate costly allowance for expansion.
Noncorrosive and nontoxic properties, eliminating the necessity for special alloys and costly measures to protect the operator’s safety.
Low cost, to reduce initial and replacement costs.
5.22.9 Effect of Regeneration Gas on Outlet Gas Quality
Regeneration gas desorbs molecular sieve beds chromatographically in the reserve order of the adsorption bead. For example:
Adsorbed methane and ethane would be desorbed first, then propanes and heavier hydrocarbons, then carbon dioxide, followed by any hydrogen sulfide that might have been in the inlet gas, and last of all, the water. The effect of the concentration of these impurities in the regeneration gas stream may be significant when regeneration gas is 10 to 15% of the net inlet gas.
In the regeneration circuit, the bulk of the water and some heavy hydrocarbons are removed from the system:
They may render the sales gas off specification for a short period.
The peak of ethane could cause the sales gas to exceed its heating value.
Concentrations of 3 to 4 ppm of H2S can be concentrated up to 20 times that amount, and thus render the composite stream far off spec.
Figure 5-49 shows the cooled regeneration gas stream is recombined with the main gas inlet to be processed. This recycle essentially eliminates the problem of making the sales gas off-specification, but it adds cost to the extent that the main gas processing capacity must be increased appropriately.
If the sales gas limits are no problem, or if there is other downstream processing, the cooled, scrubbed regeneration gas may be admitted directly to the dried outlet gas without this recycle.
5.22.10 Pressure Drop Considerations
To achieve acceptable dehydration and extend the life of the desiccant, the pressure drop through the dehydration tower should not exceed 8 psi.
The combination of feed rate, pressure drop, and adsorbent crush strength dictates the adsorption bed geometry. As noted in the above discussion regarding minimizing MTZ thickness, the bed diameter should be kept small. This feature also reduces the wall thickness of the high-pressure vessels and increases the superficial velocity, which improves mass transfer in the gas phase. However, it does not affect intra-particle mass transfer, which is the slower of the two processes.
Trent (2004) presents data that show a change in the L/D from 0.8 to 2.7 in the bed increases the useful adsorption capacity from 8.7 to 10.0 wt% in useful water capacity for an equal amount of gas dried. However, the pressure drop increases from 0.4 to 4.3 psi.
Pressure drop through the tower can be estimated from either desiccant pressure drop curves furnished by the manufacturer (Figure 5-53), or Equation 5-17 which will be followerd.
Fig. 5-53. Typical pressure-drop curve for silica gel type desiccants, 0.15-inch diameter beads.
5.22.11 Equipment
The proper selection of equipment is essential to good operations.
5.22.11.1 Inlet Gas Cleaning Equipment
All hydrocarbon liquids, free water, glycol, amine, or lube oil carry over must be cleaned from the inlet gas to ensure the best dry desiccant dehydrator operation. In all cases, the dry bed unit should have a scrubber (or a filter separator) between it and a primary well fluid separator. A microfiber filter separator (or its equivalent) should always be installed upstream of the inlet scrubber if a carryover of glycols, amines, or compressor lube oils is possible. Liquid level controls need to be checked frequently as well as the liquid dump line to ensure their operability.
5.22.11.2 Adsorber Tower
General considerations
An adsorber is a cylindrical tower filled with a solid desiccant. The depth of desiccant will vary from a few feet to 30 feet or more.
On top of the bed, a hold-down screen is provided, again covered with a layer of ceramic balls. In some cases, a layer of less expensive desiccant can be installed on the top of the bed to catch contaminants such as free water, glycol, hydrocarbons, amines, etc. This may extend the bed life. Good inlet separation of entrained contaminants is absolutely essential for long desiccant life.
Vessel diameter may be as much as 10 to 15 feet or more.
Bed height to diameter (L/D) ratio of 2.5–4.0 is desirable.
Lower ratios (1:1) are sometimes used, which could result in poor gas dehydration caused by:
Non-uniform flow
Channeling
Inadequate contact time between the wet gas and the desiccant
Three problems that frequently cause poor operation are:
Insufficient gas distribution
Inadequate insulation
Improper bed supports
1. Insufficient gas distribution
Poor gas distribution at the inlet and outlet of the desiccant beds has caused many costly problems, resulting in:
Channeling
Desiccant damage
The inlet gas distributor should be provided with adequate baffling before the gas enters the desiccant bed. Neither gas to be dehydrated nor the regeneration gas should impinge directly on the bed.
Channeling, high localized velocities and swirling can cause:
Desiccant attrition
High-pressure drop through the desiccant bed as attrition fines lodge between the regular particles
Screen-wrapped slotted pipe, with gas at low velocities exiting radially into the vessel is recommended. A 4- to 6-inch layer of large diameter (2 inch) support balls can be placed on top of the desiccant bed to improves gas distribution and prevents desiccant damage from swirling.
Swirling can destroy several feet of castable refractory lining by turning the powdered desiccant into a sandblasting agent which results in high heat losses and poor desiccant regeneration.
2- Inadequate Insulation
Internal or external insulation can be used.
Internal insulation, reduces the total regeneration gas requirements and costs.
Provision must be made for expansion and contraction so that there will be no cracking or weld failures. The lining is normally made from a castable refractory lining. Liner cracks permit some of the wet gas to bypass the desiccant bed where, a small amount of wet, bypass gas can cause freeze up in cryogenic plants. Ledges installed every few feet along the vessel wall can help eliminate liner cracks.
3- Improper Bed Supports
Two common bed supports include:
Horizontal screen supported by I-beams and a welding ring
Vessel whose bottom head is filled with graduated support balls
Screens are usually made of stainless steel or monel that have openings at least 10 meshes smaller than the smallest desiccant particle.
Screens should be securely fastened in the vessel. Provisions should be made for expansion and contraction as the adsorbers heat and cool. Annular space between the vessel wall and the edge of the bed support screen must be sealed to prevent the loss of desiccant:
Asbestos rope packing, forced in this space, is used.
A support ring around the edges of the screen is beneficial.
Support balls on the screens are helpful. 2 to 3 inches of ½-inch balls are gently placed on the screen and a 2 or 3 inch smooth layer of 1/4-inch balls is gently placed on top of the ½-inch balls. Bottom bed support typically includes three to five layers of inert ceramic balls in graduated sizes (smallest on top).
These layers prevent desiccant dust or whole particles from plugging the screen openings and forcing a high-pressure drop across the desiccant beds.
When calculating the regeneration needs of the system, it is important to include the heat requirements for the support balls.
If the bottom head of the vessel is filled with graduated support balls, a gas distributor may be required between the balls and the lower portion of the desiccant bed when upflow heating or cooling is used. This is important on large-diameter vessels to prevent channeling and poor reactivation of the desiccant. Many adsorbers have a void area in the bottom, below the bed supports, to collect contaminants, dust, and fines. A blowdown nozzle can be provided to discharge these materials. A moisture sample probe should be located in the adsorbers in cryogenic plants several feet from the outlet end of the bed and extending to the center.
This probe, used in conjunction with the outlet gas moisture probe, offers valuable flexibility in studying and solving dehydrator problems, particularly for determining if gas is being channeled down the walls of the vessel. It permits capacity tests for optimizing drying cycle times.
Since solid desiccants can produce dust, 1µm filters are frequently installed at the outlet of the dehydration unit to protect downstream equipment.
Pressurization
For best performance and maintenance of desiccant quality, adsorbers should:
Never be pressurized faster than 50 psi/min
Never be depressurized faster than 10 psi/min
Downflow pressure drop should not exceed 1 psi/ft.
Upflow pressure drop should not be less than 1/4 psi/ft.
Even with the best designs, some desiccant dust is swept out of the beds at design gas-flow rates. Certain amounts can be tolerated in many field dehydration systems.
It is not acceptable in turbo expander plant designs that involve extensive downstream heat exchange and processing. The problem is particularly significant where plate-fin or core-type heat exchangers are used. In many instances, this problem can be solved with microfiber filters (cleaning to 1 micron) with a differential pressure across them of 15 psi.
Fig. 5-54. Molecular sieve gas dehydration tower
5.22.11.3 Regeneration Gas Exchangers, Heaters, and Coolers
A gas exchanger is usually designed with the following assumptions:
All of the water will be liberated from the bed in 1 hour at 250 0F.
Regeneration gas can be cooled to within 10 0F of the sales gas temperature.
A regenerative gas heater is sized to provide:
Heat to desorb the water
Heat for the desiccant of between 500 0 and 550 0F
Heat the contactor shell
5.22 Effect of Process Variables
For the commonly used 4A molecular sieve, the Engineering Data Book suggests that the design water content of a molecular sieve when at equilibrium with saturated gas at 75°F (24°C) will be 13 lb H2O/100 lb sieve compared with a new molecular sieve, which holds about 20 lb H2O/100 lb sieve. Two main factors affect this number: water content of entering gas and adsorption temperature.
5.22.1 Quality of Inlet Gas
Performance of dry bed dehydrator is affected by moisture content and components in the natural gas stream. The relative saturation of the inlet gas determines the size of a given desiccant bed, and affects the transfer of water to the adsorbent.
If the gas comes to the dehydration unit fully saturated, which is often the case, cooling the gas and removing the condensed water before the gas enters the bed lowers water loading.
Compounds in produced natural gas adversely affect performance of the dry bed dehydrator.
Components of concern are carbon dioxide, heavy hydrocarbons, and sulfur-bearing compounds
The greater the molecular weight of a compound, the greater its adsorption potential.
Trace amounts of oxygen affect bed life and performance in a variety of ways. At the normal regeneration gas temperature of 600°F for molecular sieves, 2 moles of oxygen react with methane to form 2 moles of water and 1 mole of CO2. As this reaction is exothermic, higher amounts of oxygen in the gas can lead to temperatures above the design temperature of the molecular sieve vessel. When oxygen is present, the temperature of the beds during regeneration must be monitored for safety reasons. Oxygen undergoes partial oxidation reactions with heavier hydrocarbons, which are adsorbed in the binder, to form alcohols and carbocylic acids that ultimately turn to water and CO2. If H2S is present, it undergoes oxidation to elemental sulfur, sulfur dioxide, and water. Oxygen concentrations greater than 20 ppmv also generate olefins that become coke in the bed. These reactions reduce molecular sieve capacity by forming solid deposits and by causing incomplete removal of water during regeneration because the partial pressure of water is higher (see Figure 5.47).
To avoid the above reactions as well as COS formation, regeneration temperatures are lowered to the 300 to 375°F range. However, this range increases the required regeneration time and the amount of regeneration gas used, which increases recompression cost.
As in all processes, ensuring that the beds are protected from entrained water and hydrocarbons is important. Even trace amounts of entrained water load the bed quickly and increase the regeneration heat load.
5.22.2 Temperature
Operation is very sensitive to the temperature of the incoming gas.
Adsorption efficiency decreases as the temperature increases.
Molecular sieves and most other adsorbents have significantly higher adsorptive capacity at low temperatures. Cooling the gas to lower temperature in order to improve adsorption is limited due to condensation and/or hydrate formation.
Temperature of the regeneration gas that commingles with the incoming wet gas ahead of the dehydrators is important. The temperature must remain within 10 0F to 15 0F, otherwise liquid water and hydrocarbons will condense as the hotter gas stream cools. Condensed liquids that strike the bed can shorten the solid desiccant’s life.
Temperature of the hot gas entering and leaving a desiccant tower during the heating cycle affects plant efficiency and the desiccant life. High regeneration gas temperature assures good removal (desorption) of water and contaminants from the bed. 450 0F to 600 0F is usually used as a regeneration temperature.
Desiccant bed temperature reached during the cooling cycle is important.
If wet gas is used to cool the desiccant: Terminate the cooling cycle when the bed reaches 125 0F. Additional cooling may cause water to be adsorbed from the wet gas stream and preload (presaturate) the bed before the next adsorption cycle begins. If dry gas is used to cool the desiccant: Terminate the cooling cycle within 10 0F to 20 0F of the incoming gas temperature. It maximizes adsorption capacity of the bed.
The temperature of the regeneration gas going through the regeneration gas scrubber should be held low enough to condense and remove the water and hydrocarbons without causing hydrate problems.
5.22.3 Pressure
The adsorption capacity of a dry bed unit decreases as pressure is lowered and with usage.
Operating dry bed dehydrators well below the design pressure requires the desiccant to work harder to remove the additional water, and maintain the desired effluent dew point. With the same volume of incoming gas, the increased gas velocity occurring at the lower pressure could affect the effluent moisture content, and damage the desiccant.
At pressure above 1300 to 1400 psia, the co-adsorption effects of hydrocarbons are very significant.
5.22.4 Cycle Time
In principle, beds can be run until the first sign of breakthrough. This practice maximizes cycle time, which extends bed life because temperature cycling is a major source of bed degeneration, and minimizes regeneration costs. However, most plants operate on a set time cycle to ensure no adsorbate breakthrough. Adsorbent capacity is not a fixed value and declines with usage. For the first few months of operation, a new desiccant normally has a high capacity for water removal. If a moisture analyzer is used on the effluent gas, a much longer drying cycle can be achieved. As the desiccant ages, the cycle time can be shortened to save regeneration fuel costs and improve the desiccant life. Common cycle times are as follows: 8 hours on stream - 5 to 6 hours heating - 2 to 3 hours cooling
5.22.5 Gas Velocities
As the gas velocity during the drying cycle decreases, the ability of the desiccant to dehydrate the gas increases. On the surface, it would seem desirable to operate at minimum flow rates to utilize the desiccant fully.
Low linear velocities:
Require towers with large cross-sectional areas to handle a given gas flow
Allow wet gas to channel through the desiccant bed and thus not be properly dehydrated.
Compromise must be made between the tower diameter and the maximum utilization of the desiccant.
High linear velocities:
Lower adsorption efficiency
May cause desiccant damage
Higher inlet compression discharge pressures to maintain the same refrigeration requirements and outlet pressure.
Increased mechanical load on the adsorbent, which leads to particle breakdown and causes further increases in pressure drop.
5.22.6 Source of Regeneration Gas
Source of regeneration gas depends on plant requirements and the availability of a suitable gas stream. Regeneration gas should be dry when low effluent moisture contents (in the range of 0.1 ppm) are required. Plant tail gate gas can normally be used.
If only moderate drying is required, a portion of the wet feed gas can be used.
Figure 5-52 is an equilibrium diagram showing lines of constant water loading. For example:
A molecular sieve bed at 100 0F in equilibrium with a gas having a -80 0F water dew point will contain about 4 wt.% water. Equilibrium curves for a given adsorbate-adsorbent can be used to estimate the regeneration conditions necessary to provide the required outlet conditions. For example, If the regeneration gas is taken from inlet gas with a dew point of 40 0F and is heated to 450 0F, the mol sieve will contain 3 wt.% water after regeneration.
If the gas to be treated is at 100 0F, the intersection of the 3 wt.% line with an adsorbent temperature of 100 0F gives the minimum attainable dew point at -95 0F.
If this dew point is not satisfactory, either the regeneration gas must be heated to above 450 0F or a gas of a higher dew point (e.g., residue gas) must be used for regeneration gas.
Fig. 5-52 Equilibrium diagram showing lines of constant water loading for a type 4a molecular sieve.
5.22.7 Direction of Gas Flow
Direction of flow during the drying cycle is downward, which:
Permits higher velocities without lifting or fluidizing the desiccant bed (fluidization can severely damage the desiccant)
Direction of flow during the heating cycle is counter-current to the direction of the adsorption flow.
It permits better reactivation of the lower portion of the desiccant bed, which must perform the super-dehydration during the drying cycle, especially in cryogenic plants.
If flow is co-current, all water and/or other contaminants must move through the entire bed, thus causing additional desiccant contamination and requiring longer regeneration times.
Direction of flow during the cooling cycle:
When dry gas is used, the flow direction is counter current to the adsorption flow, thus simplifying piping and valve configuration.
When wet gas is used, the flow direction is in the same direction as the adsorption flow so that the water adsorbed during the cooling cycle as the desiccant cools will preload on the inlet end of the bed.
If counter current flow is used in cooling with wet gas, water is deposited on the exit end of the bed. When the next adsorption cycle begins, the wet gas is immediately dried, but as the dry gas continues to move down through the bed, it picks up some of the water deposited during the cooling cycle and sometimes puts too much moisture in the effluent stream.
If wet gas is used, the additional water load, deposited during the cooling cycle, should be included when the amount of desiccant needed for dehydration is calculated.
5.22.8 Desiccant Selection
Desiccant selection is based upon:
Economics
Process conditions
Desiccants are usually interchangeable. Equipment designed for one desiccant can often operate effectively with another. No desiccant product will remain effective with massive liquid carryovers.
All desiccants are damaged by heavy impurities carried into the bed with gases. These include:
Crude oil and condensate
Glycols and amines
Most corrosion inhibitors
Well treating fluids
All desiccants exhibit a decrease in capacity (design loading) with an increase in temperature.
Molecular sieves are less affected and Aluminas are most affected.
Aluminas and molecular sieves act as a catalyst with H2S to form COS, which deposits sulfur on the desiccant bed during regeneration. (Carbonyl sulfide (COS) is formed in the following reaction: H2S + CO2 ↔ H2O + COS. Its concentration in feed gas is normally extremely low.)
Alumina gels, activated aluminas, and molecular sieves are all chemically attacked by strong mineral acids and thus decrease their adsorptive capacity.
Table 5-7 provides certain physical characteristics of the more common solid desiccants.
5.22.8.1 Molecular Sieves
Molecular sieves are a class of aluminosilicates. They produce the lowest water dewpoints, and can be used to simultaneously sweeten and dry gases and liquids. Their equilibrium water capacity is much less dependent on adsorption temperature and relative humidity. They are usually more expensive.
Molecular sieves offer the highest adsorptive capacity of all desiccants when the feed gas is at very high temperatures or at low relative saturation. It is the only desiccants capable of dehydrating gas to less than 1 ppm of water content are required for cryogenic temperatures (lb/MMscf ~ ppmv / 21.4) (dew points down to -150 0F). Therefore, for gas going into cryogenic processing, the only adsorbent that can obtain the required dehydration is a molecular sieve. Of these, 4A is the most common, but the smaller pore 3A is sometimes used. It has the advantage of being a poorer catalyst for generation of COS if both H2S and CO2 are present because a portion of the more active sodium cations in 4A has been replaced with potassium. (Carbonyl sulfide (COS) is formed in the following reaction: H2S + CO2 ↔ H2O + COS. The equilibrium constant for the reaction is of the order of magnitude of 10−6 at adsorption temperatures but increases to 10−4 at regeneration temperatures.)
If both oxygen and H2S are present 3A reduces the production of elemental sulfur that can block adsorbent pores. However, plant operators usually have little incentive to use 3A for dehydrating gas going to hydrocarbon recovery.
5.22.8.2 Silica Gel
Silica Gel and Alumina are generally offer a lower first cost.
Silica Gel is essentially pure silicon dioxide, SiO2. It is used for gas and liquid dehydration and hydrocarbon (iC5+) recovery from natural gas. When used for hydrocarbon removal, the units are often called HRUs (Hydrocarbon Recovery Units) or SCUs (Short Cycle Units). When used for dehydration, silica gel will give outlet dewpoints of approximately –60°F.
Silica gel can be regenerated to a lower water content than molecular sieves and at much lower temperatures (400 0F for gels versus 500 0 to 600 0F for sieves).
It shatters in the presence of free water or light hydrocarbon liquids. The problem is minimized by using a 4 to 6 inch buffer bed of mullite ball (or equivalent) to protect the silica gel from direct contact. Silica gels are used mostly where a high concentration of water (>1 mol%) vapor is present in the feed, and low levels of water in the dehydrated gas are not needed. They are relatively noncatalytic compounds.
5.22.8.3 Aluminas
Alumina is a hydrated form of alumina oxide (Al2O3). It is used for gas and liquid dehydration and will give outlet dewpoints of about –90°F. Less heat is required to regenerate alumina and silica gel than for molecular sieve, and the regeneration temperature is lower. Molecular sieves give lower outlet water dewpoints. Aluminas are very polar and strongly attract water and acid gases. They are used for moderate levels of water in the feed when low levels of water in the product are not required. They have the highest mechanical strength of the adsorbents considered here.
5.22.8.4 Desirable Characteristics of Solid Desiccants
High adsorptive capacity (lb/lb), which reduces contactor size.
Easy regeneration, for simplicity and economics of operation.
High rate of adsorption, which allows higher gas velocities and thereby reduces contactor size.
Low resistance to gas flow, to minimize gas pressure drop through the unit.
High adsorptive capacity retained after repeated regeneration, allowing smaller initial charge and longer service before replacement.
High mechanical strength, to resist crushing and dust formation.
Inert chemicals, to prevent chemical reactions during adsorption and regeneration.
Volume unchanged when product is wet, which would otherwise necessitate costly allowance for expansion.
Noncorrosive and nontoxic properties, eliminating the necessity for special alloys and costly measures to protect the operator’s safety.
Low cost, to reduce initial and replacement costs.
5.22.9 Effect of Regeneration Gas on Outlet Gas Quality
Regeneration gas desorbs molecular sieve beds chromatographically in the reserve order of the adsorption bead. For example:
Adsorbed methane and ethane would be desorbed first, then propanes and heavier hydrocarbons, then carbon dioxide, followed by any hydrogen sulfide that might have been in the inlet gas, and last of all, the water. The effect of the concentration of these impurities in the regeneration gas stream may be significant when regeneration gas is 10 to 15% of the net inlet gas.
In the regeneration circuit, the bulk of the water and some heavy hydrocarbons are removed from the system:
They may render the sales gas off specification for a short period.
The peak of ethane could cause the sales gas to exceed its heating value.
Concentrations of 3 to 4 ppm of H2S can be concentrated up to 20 times that amount, and thus render the composite stream far off spec.
Figure 5-49 shows the cooled regeneration gas stream is recombined with the main gas inlet to be processed. This recycle essentially eliminates the problem of making the sales gas off-specification, but it adds cost to the extent that the main gas processing capacity must be increased appropriately.
If the sales gas limits are no problem, or if there is other downstream processing, the cooled, scrubbed regeneration gas may be admitted directly to the dried outlet gas without this recycle.
5.22.10 Pressure Drop Considerations
To achieve acceptable dehydration and extend the life of the desiccant, the pressure drop through the dehydration tower should not exceed 8 psi.
The combination of feed rate, pressure drop, and adsorbent crush strength dictates the adsorption bed geometry. As noted in the above discussion regarding minimizing MTZ thickness, the bed diameter should be kept small. This feature also reduces the wall thickness of the high-pressure vessels and increases the superficial velocity, which improves mass transfer in the gas phase. However, it does not affect intra-particle mass transfer, which is the slower of the two processes.
Trent (2004) presents data that show a change in the L/D from 0.8 to 2.7 in the bed increases the useful adsorption capacity from 8.7 to 10.0 wt% in useful water capacity for an equal amount of gas dried. However, the pressure drop increases from 0.4 to 4.3 psi.
Pressure drop through the tower can be estimated from either desiccant pressure drop curves furnished by the manufacturer (Figure 5-53), or Equation 5-17 which will be followerd.
Fig. 5-53. Typical pressure-drop curve for silica gel type desiccants, 0.15-inch diameter beads.
5.22.11 Equipment
The proper selection of equipment is essential to good operations.
5.22.11.1 Inlet Gas Cleaning Equipment
All hydrocarbon liquids, free water, glycol, amine, or lube oil carry over must be cleaned from the inlet gas to ensure the best dry desiccant dehydrator operation. In all cases, the dry bed unit should have a scrubber (or a filter separator) between it and a primary well fluid separator. A microfiber filter separator (or its equivalent) should always be installed upstream of the inlet scrubber if a carryover of glycols, amines, or compressor lube oils is possible. Liquid level controls need to be checked frequently as well as the liquid dump line to ensure their operability.
5.22.11.2 Adsorber Tower
General considerations
An adsorber is a cylindrical tower filled with a solid desiccant. The depth of desiccant will vary from a few feet to 30 feet or more.
On top of the bed, a hold-down screen is provided, again covered with a layer of ceramic balls. In some cases, a layer of less expensive desiccant can be installed on the top of the bed to catch contaminants such as free water, glycol, hydrocarbons, amines, etc. This may extend the bed life. Good inlet separation of entrained contaminants is absolutely essential for long desiccant life.
Vessel diameter may be as much as 10 to 15 feet or more.
Bed height to diameter (L/D) ratio of 2.5–4.0 is desirable.
Lower ratios (1:1) are sometimes used, which could result in poor gas dehydration caused by:
Non-uniform flow
Channeling
Inadequate contact time between the wet gas and the desiccant
Three problems that frequently cause poor operation are:
Insufficient gas distribution
Inadequate insulation
Improper bed supports
1. Insufficient gas distribution
Poor gas distribution at the inlet and outlet of the desiccant beds has caused many costly problems, resulting in:
Channeling
Desiccant damage
The inlet gas distributor should be provided with adequate baffling before the gas enters the desiccant bed. Neither gas to be dehydrated nor the regeneration gas should impinge directly on the bed.
Channeling, high localized velocities and swirling can cause:
Desiccant attrition
High-pressure drop through the desiccant bed as attrition fines lodge between the regular particles
Screen-wrapped slotted pipe, with gas at low velocities exiting radially into the vessel is recommended. A 4- to 6-inch layer of large diameter (2 inch) support balls can be placed on top of the desiccant bed to improves gas distribution and prevents desiccant damage from swirling.
Swirling can destroy several feet of castable refractory lining by turning the powdered desiccant into a sandblasting agent which results in high heat losses and poor desiccant regeneration.
2- Inadequate Insulation
Internal or external insulation can be used.
Internal insulation, reduces the total regeneration gas requirements and costs.
Provision must be made for expansion and contraction so that there will be no cracking or weld failures. The lining is normally made from a castable refractory lining. Liner cracks permit some of the wet gas to bypass the desiccant bed where, a small amount of wet, bypass gas can cause freeze up in cryogenic plants. Ledges installed every few feet along the vessel wall can help eliminate liner cracks.
3- Improper Bed Supports
Two common bed supports include:
Horizontal screen supported by I-beams and a welding ring
Vessel whose bottom head is filled with graduated support balls
Screens are usually made of stainless steel or monel that have openings at least 10 meshes smaller than the smallest desiccant particle.
Screens should be securely fastened in the vessel. Provisions should be made for expansion and contraction as the adsorbers heat and cool. Annular space between the vessel wall and the edge of the bed support screen must be sealed to prevent the loss of desiccant:
Asbestos rope packing, forced in this space, is used.
A support ring around the edges of the screen is beneficial.
Support balls on the screens are helpful. 2 to 3 inches of ½-inch balls are gently placed on the screen and a 2 or 3 inch smooth layer of 1/4-inch balls is gently placed on top of the ½-inch balls. Bottom bed support typically includes three to five layers of inert ceramic balls in graduated sizes (smallest on top).
These layers prevent desiccant dust or whole particles from plugging the screen openings and forcing a high-pressure drop across the desiccant beds.
When calculating the regeneration needs of the system, it is important to include the heat requirements for the support balls.
If the bottom head of the vessel is filled with graduated support balls, a gas distributor may be required between the balls and the lower portion of the desiccant bed when upflow heating or cooling is used. This is important on large-diameter vessels to prevent channeling and poor reactivation of the desiccant. Many adsorbers have a void area in the bottom, below the bed supports, to collect contaminants, dust, and fines. A blowdown nozzle can be provided to discharge these materials. A moisture sample probe should be located in the adsorbers in cryogenic plants several feet from the outlet end of the bed and extending to the center.
This probe, used in conjunction with the outlet gas moisture probe, offers valuable flexibility in studying and solving dehydrator problems, particularly for determining if gas is being channeled down the walls of the vessel. It permits capacity tests for optimizing drying cycle times.
Since solid desiccants can produce dust, 1µm filters are frequently installed at the outlet of the dehydration unit to protect downstream equipment.
Pressurization
For best performance and maintenance of desiccant quality, adsorbers should:
Never be pressurized faster than 50 psi/min
Never be depressurized faster than 10 psi/min
Downflow pressure drop should not exceed 1 psi/ft.
Upflow pressure drop should not be less than 1/4 psi/ft.
Even with the best designs, some desiccant dust is swept out of the beds at design gas-flow rates. Certain amounts can be tolerated in many field dehydration systems.
It is not acceptable in turbo expander plant designs that involve extensive downstream heat exchange and processing. The problem is particularly significant where plate-fin or core-type heat exchangers are used. In many instances, this problem can be solved with microfiber filters (cleaning to 1 micron) with a differential pressure across them of 15 psi.
Fig. 5-54. Molecular sieve gas dehydration tower
5.22.11.3 Regeneration Gas Exchangers, Heaters, and Coolers
A gas exchanger is usually designed with the following assumptions:
All of the water will be liberated from the bed in 1 hour at 250 0F.
Regeneration gas can be cooled to within 10 0F of the sales gas temperature.
A regenerative gas heater is sized to provide:
Heat to desorb the water
Heat for the desiccant of between 500 0 and 550 0F
Heat the contactor shell
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Basics of Gas Field Processing
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Gas Dehydration - Chapter 5 - Part 4
5.22.11.4 Regeneration Gas Separator
Most desiccants also have an affinity for hydrocarbons, thus a skimmer is used to separate the valuable hydrocarbons from the water to be discarded. Frequent pH tests on the discarded water helps pinpoint corrosion problems in the adsorption system. A common problem encountered in regeneration gas separators is the fouling of the liquid dump line by desiccant dust and heavy oils.
5.22.11.5 Expander Plant Molecular Sieve Applications
Turbo expander plants commonly operate down to temperatures of -150 0F.
Operating points much below the equilibrium water content data illustrated in McKetta-Wehe chart (include designs to water contents as low as 1 ppm).
As shown in Table 5-7 (UP), only molecular sieves and activated alumina are capable of such performance. Molecular sieves are used in approximately 95% of the dehydration equipment for this type of plant (a 4A molecular sieve has twice the adsorptive capacity of activated alumina).
5.23 Desiccant Performance
5.23.1 General Conditions
Desiccants decline in adsorptive capacity at different rates under varying operating conditions.
Desiccant aging is a function of many factors, including:
Number of cycles experienced
Exposure to any harmful contaminants present in the inlet stream.
The most important variable affecting the decline rate of desiccant capacity is the chemical composition of the gas or liquid to be dried.
Capacity of a new desiccant will decline slowly during the first few months in service because of cyclic heating, cooling, and netting. Desiccant capacity usually stabilizes at about 55 to 70% of the initial capacity.
5.23.2 Moisture Analyzer
The moisture analyzer is used to optimize the drying cycle time, where the drying time is always shortened as the desiccant ages. Both inlet and outlet moisture analyzer probes should be used. A probe extending approximately 2 feet upward into the bed from the outlet end is recommended because it allows a dehydration capacity test to be run without the risk of a water breakthrough.
5.23.3 Effect of Contaminants in Inlet Feed Stream
Compressor oils, corrosion inhibitors, glycols, amines, and other high-boiling contaminants cause a decline in desiccant capacity, because normal reactivation temperatures will not vaporize the heavy materials. Residual contaminants slowly build up on the desiccant’s surface reducing the area available for adsorption. Many corrosion inhibitors chemically attack certain desiccants, permanently destroying their usefulness.
5.23.4 Effect of Regeneration Gases Rich in Heavy Hydrocarbons
Use of this rich gas in a 5500 to 600 0F regeneration service aggravates coking problems. Rich gases may be dried satisfactorily with molecular sieves. Lean dry gas is always preferable for regeneration.
5.23.5 Effect of Methanol in the Inlet Gas Stream
Methanol in the inlet gas is a major contributor to the coking of molecular sieves where regeneration is carried out at temperatures above 550 0F. Polymerization of methanol during regeneration produces dimethyl ether and other intermediates that will cause coking of the beds.
Conversion to ethylene glycol injection, instead of methanol for hydrate control, will increase sieve life and add at least 10% to sieve capacity by removing the vapor phase methanol from the system.
5.23.6 Useful Life
The most common reasons for replacing a bed are loss of adsorbent capacity and unacceptable pressure drop, which usually occur simultaneously. Values for the loss of capacity with time vary considerably, but common values used for molecular sieves in dehydration service are a 35% capacity loss over a 3 to 5 year period or a 50% loss in approximately 1,600 cycles. Typically, a rapid loss occurs in the beginning and a gradual loss thereafter. The adsorbent decays primarily because of carbon and sulfur fouling and caking caused by instability in the clay binder. These effects occur during bed regeneration.
Increased pressure drop is usually caused by breakdown of adsorbent into finer particles and by caking at the top of the bed because of refluxing. Attrition can occur when the pressure is increased or decreased (more than the allowable pressure-changing rate) after or before regeneration.
Monitoring the pressure drop is important, as it provides a good diagnostic to bed health.
In general, adsorbent life ranges from one to four years in normal service. Longer life is possible if feed gas is kept clean. Effectiveness of regeneration plays a major role in retarding the decline of a desiccant’s adsorptive capacity and prolonging its useful life. If all the water is not removed from the desiccant during each regeneration, its usefulness will sharply decrease.
5.23.7 Effect of Insufficient Reactivation
Insufficient reactivation can occur if the regeneration gas temperature or velocity is too low.
A desiccant manufacturer will generally recommend the optimum regeneration temperature and velocity for the product. Velocity should be high enough to remove the water and other contaminants quickly, thus minimizing the amount or residual water and protect the desiccant.
5.23.8 Effect of High Reactivation Temperature
Higher reactivation temperatures remove volatile contaminants before they form coke on the desiccant, maximizes desiccant capacity and ensures minimum effluent moisture content.
Final effluent hot gas temperature should be held one or two hours to achieve effective desiccant
reactivation.
5.24 Design
5.24.1 Pressure Drop & Minimum Diameter
The first step is to determine the bed diameter, which depends on the superficial velocity. Too large a diameter will require a high regeneration gas rate to prevent channeling. Too small a diameter will cause too high a pressure drop and damage the sieve. The pressure drop is determined by a modified Ergun equation, which relates pressure drop to superficial velocity as follows:
ΔP / L = B ս V + C ρ V2 Eq 5-17
ΔP = pressure drop, psi
L = length of packed bed, ft
B & C = constants from table 5-10
µ= viscosity, cp
V = superficial vapor velocity, ft/min
ρ = density, lb/ft3
Particle Type B C
1/8" bead (4x8 mesh) 0.0560 0.0000889
1/8" extrudate 0.0722 0.000124
1/16" bead (8x12 mesh) 0.152 0.000136
1/16" extrudate 0.238 0.000210
Table. 5-10. Constants for Equation 5-6.
Fig. 5-55 was derived from Eq 5-17 by assuming a gas composition and temperature and setting the maximum allowable ΔP/L equal to 0.33 psi/ft. The design pressure drop through the bed should be about 5 psi. A design pressure drop higher than 8 psi is not recommended as the desiccant is fragile and can be crushed by the total bed weight and pressure drop forces.
Fig. 5-55. Allowable Velocity for Mole Sieve Dehydrator
Remember to check the pressure drop after the bed height has been determined. Once the allowable superficial velocity is estimated, calculate the bed minimum diameter
(i.e., Dminimum), and select the nearest standard diameter (i.e. Dselected):
Dminimum = (4q /π Vmax)0.5 Eq. 5-18
where
D = diameter, ft
q = actual gas flow rate, ft3/min
q = m/60ρ Eq. 5-19
where
m = mass flow rate, lb/hr
ρ= density, lb/ft3
Obtain the corresponding superficial velocity, Vadjusted as follows:
Vadjusted = Vmax (Dminimum/ Dselected)2 Eq. 5-20
An alternative and more exact method to calculate the maximum superficial velocity can be determined by Eq 5-21, which was derived from Eq.5-17.
Vmax = [(ΔP/L)max/(Cρ)]0.5 – [(B/C)(ս/ρ)/2] Eq 5-21
The value of (ΔP/L)max in these equations depends on the sieve type, size, and shape, but a typical value for design is 0.33 psi/ft.
5.24.2 Mass desiccant Required & Bed Length
Choose an adsorption period and calculate the mass of desiccant required. Eight to twelve hour adsorption periods are common. Periods of greater than 12 hours may be justified especially if the feed gas is not water saturated.
Long adsorption periods mean less regenerations and longer sieve life, but larger beds and additional capital investment.
During the adsorption period, the bed can be thought of as operating with three zones.
Saturation or equilibrium zone
The middle or mass transfer zone (MTZ) is where the water content of the gas is reduced from its inlet concentration to < 1 ppm. (lb/MMscf ~ ppmv / 21.4)
The bottom zone is unused desiccant and is often called the active zone. If the bed operates too long in adsorption, the mass transfer zone begins to move out the bottom of the bed causing a “breakthrough.”
In the saturation zone, molecular sieve is expected to hold approximately 13 pounds of water per 100 pounds of sieve. New sieve will have an equilibrium capacity near 20%; 13% represents the approximate capacity of a 3-5 year old sieve.
This capacity needs to be adjusted when the gas is not water saturated or the temperature is above 75°F. Figures. 5-56 and 5.57 are used to find the correction factors for molecular sieve.
Alternatively, both parameters may be calculated using the following corellations:
CSS = 0.636 + 0.0826 ln(Sat) Eq. 5-22
and
CT = 1.20 – 0.0026 t(°F) Eq. 5-23
CT = 1.11 – 0.0047 t(°C) Eq. 5-24
where CSS and CT are correction factors for subsaturation and temperature, respectively.
Sat is the percent of saturation.
To determine the mass of desiccant required in the saturation zone, calculate the amount of water to be removed during the cycle and divide by the effective capacity.
SS = Wr / [(0.13)(CSS)(CT)] Eq 5-25
LS = 4 SS /[π (D2) (bulk density) Eq 5-26
Where
LS = length of packed bed saturation zone, ft
SS = amount molecular sieve required in saturation zone, lb
Wr = water removed per cycle, lb
CSS = saturation correction factor for sieve (Fig.5-56)
CT = temperature correction factor (Fig.5-57)
LS = length of packed bed saturation zone, ft
D = diameter, ft
Molecular sieve bulk density is 42 to 46 lb/ft3 for spherical particles and 40 to 44 lb/ft3 for extruded cylinders.
Even though the MTZ will contain some water (approximately 50% of the equilibrium capacity), the saturation zone is estimated assuming it will contain all the water to be removed.
The length of the mass transfer zone can be estimated as follows:
LMTZ = (Vadjusted/35)0.3 (ZL) Eq 5-27
Where:
LMTZ = length of packed bed mass transfer zone, ft
ZL = 1.70 ft for 1/8 inch sieve
= 0.85 ft for 1/16 inch sieve
Fig. 5-56. Mole Sieve Capacity Correction for gas percent saturation
Fig. 5-57. Mole Sieve Capacity Correction for Temperature
The total bed height is the summation of the saturation zone and the mass transfer zone heights. It should be no less than the vessel inside diameter, or 6 feet, whichever is greater.
Now the total bed pressure drop is checked. The ΔP/L for the selected diameter, Dselected, is adjusted using Eq 5-17 or the following approximation:
(ΔP / L)adjusted ≅ (0.33 psi/ft) (Vadjusted/Vmax)2 Eq 5-28
The result is multiplied times the total bed height (LS +LMTZ) to get the total design pressure drop, which should be 5- 8 psi. This is important, because the operating pressure drop can increase to as much as double the design value over three years. Too high a pressure drop plus the bed weight can crush the sieve. If the design pressure drop exceeds 8 psi, the bed diameter should be increased and the sieve amount and vessel dimensions recalculated.
To estimate the total cylindrical length of a tower, add 3 feet to the bed height, which provides the space for an inlet distributor and for bed support and hold-down balls under and on top of the sieve bed.
5.24.3 Regeneration Calculations
The first step is to calculate the total heat required to desorb the water and heat the desiccant and vessel. A 10% heat loss is assumed.
Qw = (1800 Btu/lb) (lbs of water on bed) Eq 5-29
Qsi = (lbs of sieve)(0.24 Btu/lb 0F) (Trg – Ti) Eq 5-30
Qst = (lbs of steel)(0.12 Btu/lb 0F) (Trg – Ti) Eq 5-31
Qhl =heat loss = (Qw +Qsi+Qst) (0.1) Eq 5-32
where
Qw = desorption of water heat duty, Btu
Qsi = duty required to heat mole sieve to regeneration temperature, Btu
Qst = duty required to heat vessel and piping to regeneration temperature, Btu
Qhl = regeneration heat loss duty, Btu
Trg = regeneration temperature, 0F
Ti = initial, temperature, 0F
The temperature, Trg, is the temperature to which the bed and vessel must be heated based on the vessel being externally insulated (i.e., no internal insulation which is usually the case). This is about 50°F below the temperature of the hot regeneration gas entering the tower.
The weight of the vessel steel is estimated from equations 5-33 and 5-34.
The design pressure, Pdesign, is usually set at 110% of the maximum operating pressure. The value of 0.125 in Eq 5-34 is the corrosion allowance in inches. The term 0.75Dbed is to account for the weight of the tower heads. The value of "3" provides the space for the inlet distributor and support and hold-down balls.
t(inches) = (12DbedPdesign) / (37,600 – 1.2Pdesign) Eq 5-33
Weight of steel (lb) = 155 (t + 0.125) (LS + LMTZ + 0.75Dbed + 3)Dbed Eq 5-34
For determination of the regeneration gas rate, calculate the total regeneration load from Eq. 5-35
Qtr = (2.5)(Qw + Qsi + Qst + Qhl) Eq. 5-35
where
Qtr = total regeneration heat duty, Btu
The 2.5 factor corrects for the change in temperature difference (in – out) across the bed with time during the regeneration period. It assumes that 40% of the heat in the gas transfers to the bed, vessel steel, and heat loss to atmosphere; and the balance leaves with the hot gas.
The regeneration-gas flow rate is calculated from Eq 5-36 below. The heat capacity, Cp, is calculated with Eq 5-37, with the enthalpies obtained from the enthalpy vs. temperature plots for various pressures in (GPSA- Section 24 – or refer to definition of (Cp) in eq. 5-37). The temperature, Thot, is 50°F above the temperature, Trg, to which the bed must be heated. The temperature, Tb, is the bed temperature at the beginning of regeneration, which is the same as the dehydration-plant feed temperature.
The heating time is usually 50% to 60% of the total regeneration time which must include a cooling period. Figure 5-59 shows a typical temperature profile for a regeneration period (heating and cooling). For 8 hour adsorption periods, the regeneration normally consists of 4 1/2 hours of heating, 3 hours of cooling and 1/2 hour for standby and switching. For longer periods the heating time can be lengthened as long as a minimum pressure drop of 0.01 psi/ft is maintained to ensure even flow distribution across the bed.
mrg = Qtr/[Cp(Thot -Tb)(heating time)] Eq 5-36
Cp (Btu/lb/°F) = (Hhot - Hi)/(Thot -Tb) Eq 5-37
Where
mrg = regeneration mass flow rate, lb/hr
Thot = Hot gas temperature, 0F
Tb = bed starting temperature, 0F
Cp = gas heat capacity, Btu/(lb . °F) = (Hhot - Hi)/(Thot -Tb) (Extract the value from fig. 5-60 ,which include values at 600 psia, or (Use Cp = 0.66 as an approximate value for natural gas in regeneration operation ranges) Hhot = enthalpy, BTU/lb, and Hi = enthalpy, BTU/lb. (Other values can be extracted from curves in GPSA, chapter 24 “Total Enthalpy of Paraffin Hydrocarbon Vapor”).
The superficial velocity of the regeneration gas is calculated from Eq 5-38 for which (q) is calculated from Eq 5-19
V = 4q / (π D2) Eq 5-38
where
V = Velocity, ft/sec
D = diameter, ft
q = actual gas flow rate, ft3/min
The calculated superficial velocity can not be less than the value that corresponds with a minimum bed pressure drop of 0.01 psi/ft. This can be determined from Fig. 5-58, which was derived from Eq 5-17 by assuming a gas composition and temperature and setting ΔP/L equal to 0.01 psi / ft.
If the calculated velocity is less than this, the regeneration gas rate, (mrg) must be increased by multiplying it by the ratio Vmin / V, and the period of regeneration should be decreased by multiplying it times the ratio V / Vmin. A more exact method for calculating the minimum superficial velocity is to use Eq 5-21, but to consider it in terms of (ΔP/L)min and Vmin instead of (ΔP/L)max and Vmax.
Fig. 5-58. Minimum Regeneration Velocity for Mole Sieve Dehydrator
Fig. 5-59. Inlet and Outlet Temperatures During Typical Solid Desiccant Bed Regeneration Period
5.24.4 Design Example
100 MMscfd of natural gas with a molecular weight of 18 is water saturated at 600 psia and 100°F and must be dried to –150°F dew point.
Determine the water content of the gas, and the amount of water that must be removed; and do a preliminary design of a molecular-sieve dehydration system consisting of two towers (one dehydration and the other regeneration in a cycle basis). Use 4A molecular sieve of 1/8" beads (i.e., 4x8 mesh). Compressibility factor z = 0.93.
The regeneration gas is part of the plant’s residue gas, which is at 600 psia and 100°F and has a molecular weight of 17. The bed must be heated to 500°F for regeneration. Compressibility factor = 0.95.
Solution Steps
1. Determine the bed diameter and the corresponding ΔP/L and V. First determine the maximum superficial velocity from Eq 5-10. Let the maximum ΔP/L be 0.33 psi/ft.
Vmax = [(ΔP/L)max/(Cρ)]0.5 – [(B/C)(ս/ρ)/2] Eq 5-21
Fro chapter 1, Eq. 1-9
ρg= 0.093 ((MW)P)/TZ lb/ft3 (Eq. 1-19)
ρg = 0.093 x 18 x 600 / (560 x 0.93)
ρ = 1.93 lb/ft3
Or,
ρ = density of gas (ρ)= (18 mole weight) (600 psia) / [10.73 (560 °R)(0.93 z)] = 1.93 lb/ft3
µ = 0.015 centipoise (Fig. 1-11)
C = 0.000089 from table 5-10.
Vmax = {(0.33 psi/ft) / (0.0000889)(1.93) lb/ft3)}0.5 – [(0.056 / 0.0000889) (0.015 centipoise/1.93 lb/ft3)/2]
= 41.4 ft/min
Mass flow rate (m) = (100 X 106 scf/day) (18 lb/lb mole) / [(24 hr/day)(379.5 scf/lb mole)]
m = 198,000 lb/hr of wet gas
From Eq. 5-19
q = (198,000 lb/hr)/((60 min/hr)(1.93 lb/ft3)) = 1710 ft3/min of wet gas
From eq. 5-18
Dminimum = [(4(1710 ft3/min))/(3.14 X 41.4 ft/min)]0.5 = 7.25 ft
Round off upward to 7.5 ft diameter, for which V and ΔP/L are adjusted as follows (eq. 5-20):
Vadjusted = (41.4)(7.25/7.5)2 = 38.7 ft/min
(ΔP/L)adjusted = 0.33(38.7/41.4)2 = 0.29 psi/ft
2. Estimate the amount of water to be removed from the feed per cycle for each bed.
Base this on a 24-hour cycle consisting of 12 hours adsorbing and 12 hours regenerating (heating, cooling, standby, and valve switching).
From chapter 4, fig. 4-8, the water content at 600 psia and 100°F is 88 pounds/MMscf. The water content at a dew point of –150°F is essentially zero, so the water removed is the following:
w = (88-0 lb/MMscf)(100 MMscf/day) / (24 hr/day)
= 367 lb/hr of water removed
Wr = (367 lb/hr) (12 hr) = 4404 lb water removed per 12-hour drying period or 24-hour cycle per bed.
3. Determine the amount of sieve required and the bed height based on a sieve bulk density of 45 lb/ft3 (table 5-7). Since the feed gas is water saturated, the relative humidity is 100%, so CSS is 1.0, and CT is 0.93 at 100°F from Figures. 5-56 & 5-57.
Applying the equations 5-25:
SS = Wr / [(0.13)(CSS)(CT)] Eq 5-25
SS = (4404) / ((0.13)(1.0)(0.93)) = 36,427 lb of sieve for each bed.
From eq. 5-26.
LS = 4 SS /[π (D2) (bulk density) Eq 5-26
LS = (4)(36,427) /((3.1416) (7.5)2 (45)) = 18.3 ft bed height
from eq-5-27
LMTZ = (Vadjusted/35)0.3 (ZL) Eq 5-27
LMTZ = (38.7/35)0.3 (1.7) = 1.8 ft for mass-transfer zone
LS + LMTZ = 18.3 + 1.8 = 20.1 ft of sieve for each bed
The total sieve = (20.1/18.3)(36,427) = 40,010 lb for each bed
4. Check the bed design and pressure drop which is the ΔP/L calculated in Step 1 times the total bed height calculated in Step 3:
(0.29 psi/ft)(20.1 ft) = 5.8 psi which meets the criterion of not exceeding 8 psi.
5. Calculate the total heat required to desorb the water based on heating the bed and vessel to 500°F. First calculate the weight of steel from Eq 5-33 and 5-34. Let the design pressure, Pdesign, be 110% of the operating pressure: Pdesign = (600)(1.1) = 660 psia.
t(inches) = (12DbedPdesign) / (37,600 – 1.2Pdesign) Eq 5-33
Weight of steel (lb) = 155 (t + 0.125) (LS + LMTZ + 0.75Dbed + 3)Dbed Eq 5-34
t = (12)(7.5)(660) / (37,600 – (1.2)(660)) = 1.614 inches
Weight of steel = (155) (1.614 + 0.125) (18.3 + 1.+ (0.75) (7.5) + 3) (7.5) = 58,070 pounds
From the following equations:
Qw = (1800 Btu/lb) (lbs of water on bed) Eq 5-29
Qsi = (lbs of sieve)(0.24 Btu/lb 0F) (Trg – Ti) Eq 5-30
Qst = (lbs of steel)(0.12 Btu/lb 0F) (Trg – Ti) Eq 5-31
Qhl =heat loss = (Qw +Qsi+Qst) (0.1) Eq 5-32
Qtr = (2.5)(Qw + Qsi + Qst + Qhl) Eq. 5-35
Qw = (1800 Btu/lb (4404 lb water) = 7,927,000 Btu
Qsi = (40,010 lb) (0.24 Btu/lb/°F) (500°F – 100°F) = 3,841,000 Btu
Qst = (58,070 lb) (0.12 Btu/lb/°F) (500°F – 100°F) = 2,787,000 Btu
Qhl = (2,787,000 + 7,927,000 + 3,841,000) (0.10) = 1,455,000 Btu
Qtr = (2.5) (2,787,000 + 7,927,000 + 3,841,000 + 1,455,000) = 40,025,000 Btu
6. Calculate the flow rate of regeneration gas using Eq 5-36. Let the heating time be 60% of the total regeneration period, and calculate the gas heat capacity, use Cp, = 0.66 (Btu/lb/°F), or (calculate it from Eq 5-37 using enthalpy curves in GPSA, chapter 24 “Total Enthalpy of Paraffin Hydrocarbon Vapor”):
mrg = Qtr/[Cp(Thot -Tb)(heating time)] Eq 5-36
Cp (Btu/lb/°F) = (Hhot - Hi)/(Thot -Tb) Eq 5-37
(60%) (12 hr) = 7.2 hours heating
Cp (@ 600 psia from fig 5-60) = ((545 – 250) (Btu/lb))/((550 – 100) (°F)) = 0.66 Btu/lb/°F
mrg = (40,025,000 Btu)/((0.66 Btu/lb/°F) (550 – 100) (°F) (7.2 hr)) = 18,717 lb/hr (Eq 5-22)
7. Check that the ΔP/L ≥ 0.01 psi/ft at 550°F.
ρg = 0.093 x 17 x 600 / (1110 x 0.95)
ρ = 0.9 lb/ft3
From Eq. 5-19 q = m/60ρ
= 18,717 / (60 x 0.9) = 346.6 ft3/min of hot regeneration gas
Rearranging Eq 5-38:
V = 4q / (π D2) Eq 5-38
V = 4q/πD2 = ((4)(346.6)/((3.414)(7.5)2) = 7.21 ft/min
μ = 0.023 cP (Fig. 1-11)
From ΔP / L = B ս V + C ρ V2 Eq 5-17
ΔP/L = (0.056) (0.023) (7.21) + (0.0000889) (0.9) (7.21)2 = 0.013 psi/ft (Eq 5-6)
This is safely above the minimum value of 0.01 psi/ft needed to prevent channeling.
8. The design results are summarized as follows:
Number of vessels: two
Vessel design pressure and temperature: 660 psig and 600°F
Vessel dimensions: 90 inches (7.5 feet) ID by 23.1 feet tan to tan
Weight of molecular sieve: 2x40,010 lb
Regeneration gas rate: 18,717 lb/hr (10.026 MMscfd)
Regeneration gas temperature: 550°F
Cycle time: 24 hours, 12 hours adsorption, 12 hours regeneration
[img]ttp://oilprocessing.net/data/documents/V5-60.png[/img]
Fig.5-60. Total Enthalpy of Paraffin Hydrocarbon Vapor @ 600 psia.
5.25 Nonregenerable Dehydrator
In some situations, such as remote gas wells, use of a consumable salt desiccant, such as CaCl2, may be economically feasible.
5.25.1 Calcium Chloride Dehydrator Unit
Calcium chloride (CaCI2) dehydrator consists of three sections (Figure 5-61):
Inlet gas scrubber
Brine tray
Solid brine particles
5.25.2 Principles of Operation
Solid desiccant is placed in the top of the unit. Water-wet gas contacting the solid CaCI2 gives up part of its water to form liquid brine to drip down and fill the trays.
Solid anhydrous CaCl2 combines with water to form various CaCl2 hydrates (CaCl2 .XH2O). As water absorption continues, CaCl2 is converted to successively higher states of hydration eventually forming a CaCl2 brine solution.
Inlet gas coming up through the specially designed nozzles on the trays contact the brine efficiently. The wettest gas contacts the most dilute brine (about 1.2 specific gravity).
Approximate 2.5 Lb H2O/lb CaCI2 is removed in the trays. Brine gravity on the top tray is about 1.4. Another 1 lb H2O/lb CaCI2 is removed in the solid bed section.
Outlet water contents of 1 lb/MMscf have been achieved with CaCl2 dehydrators. Typical CaCl2
Maximum dew point depression of 60 0 to 70 0F occurs in this section. Typically used in remote, small gas fields without heat or fuel. capacity is 0.3 lb CaCl2 per lb H20. Superficial bed velocities are 20-30 ft/min and length to diameter ratio for the bed should be at least 3 to 4:1.
CaCl2 dehydrators may offer a viable alternative to glycol units on low rate, remote dry gas wells. The CaCl2 must be changed out periodically. In low capacity – high rate units this may be as often as every 2-3 weeks. Brine disposal raises environmental issues. In addition, under certain conditions the CaCl2 pellets can bond together to form a solid bridge in the fixed bed portion of the tower. This results in gas channeling and poor unit performance.
Advantages Disadvantages
Simple
No moving parts
No heat required
Does not react with H2S or CO2
Can dehydrate hydrocarbon liquids Batch process
Emulsifies with oil
Unreliable
Limited dew point depression
Table. 5-11. Advantages and disadvantage of Calcium Chloride Dehydrator Unit
Operating Problems
Bridging and channeling is a problem.
Brine can crystallize at 85 0F, thus during low flow periods can plug vessel outlet or trays.
Brine carry-over can cause severe corrosion problems.
Fig. 5-61. CaCl2 dehydrator.
Design Considerations
Figure 5-62 illustrates the water content of natural gas dried by solid calcium chloride bed units.
Fig. 5-62 Water content on natural gas driven by CaCl2 unit (Left: freshly recharged; right: just prior to recharging).
5.26 Dehydration by Refrigeration
The dehydration of natural gas can also be achieved by refrigeration and/or cryogenic processing down to – 150°F in the presence of methanol hydrate and freeze protection. The condensed water and methanol streams decanted in the cold process can be regenerated by conventional distillation or by a patented process called IFPEX-1®.
In the latter process illustrated in schematic form in Figure 5-63 a slip stream of water saturated feed gas strips essentially all the methanol in the cold decanted methanol water stream originating in the cold process at feed gas conditions to recirculate the methanol to the cold process. The water stream leaving the stripper contains generally less than 100 ppm wt of methanol. No heat is required for the process and no atmospheric venting takes place.
The process has several major advantages:
• It can obtain dew points in the −100 to −150°F (−70 to –100°C) range.
• It requires no heat input other than to the methanol regenerator.
• It requires no venting of hydrocarbon-containing vapors.
However, it requires external refrigeration to cool the gas, and minimal methanol losses occur in the stripper.
Fig. 5-63. Example IFPEX-1 . Dehydration Process Flow Diagram
5.27 Dehydration by Membrane Permeation
Membranes can be used to separate gas stream components in natural gas such as water, CO2 and hydrocarbons according to their permeabilities. Each gas component entering the separator has a characteristic permeation rate that is a function of its ability to dissolve in and diffuse through the membrane.
The driving force for separation of a gas component in a mixture is the difference between its partial pressure across the membrane. As pressurized feed gas flows into the metal shell of the separator, the fast gas component, such as water and CO2, permeate through the membrane. This permeate is collected at a reduced pressure, while the non-permeate stream, i.e., the dry natural gas, leaves the separator at a slightly lower pressure than the feed.
The amount of methane and other natural gas components in the permeate stream is dependent on pressure drop and the surface area of the membranes. However, 5–10% of the feed stream is a realistic figure. Dehydration by membrane permeation is therefore normally only considered for plants that can make use of low pressure natural gas fuel.
Membranes are characteristic by lightweight, large turndown ratio, and low maintenance, that make them competitive with glycol units in some situations.
Feed pretreatment is a critical component of a membrane process. The inlet gas must be free of solids and droplets larger than 3 microns.
Inlet gas temperature should be at least 20°F (10°C) above the dew point of water to avoid condensation in the membrane.
Units operate at pressures up to 700 to 1,000 psig (50 – 70 barg) with feed gases containing 500 to 2,000 ppmv of water (lb/MMscf ~ ppmv / 21.4). They produce a product gas stream of 20 to 100 ppmv and 700 to 990 psig (48 to 68 barg). The low-pressure (7 to 60 psig [0.5 to 4 barg]) permeate gas volume is about 3 to 5% of the feed gas volume.
This gas must be recompressed or used in a low-pressure system such as fuel gas.
Smith (2004) suggests that membranes used for natural gas dehydration are economically viable only when dehydration is combined with acid-gas removal.
On the basis of commercial units installed and several studies; (Bikin et al., 2003), membranes are economically attractive for dehydration of gas when flow rates are less than 10 MMscfd (0.3 MMSm3/d). Binci et al. claim that membrane units are competitive with TEG dehydrators on offshore platforms at flows below 56 MMscfd (1.6 MMSm3/d). Certainly, the reliability and simplicity of membranes make them attractive for offshore and remote-site applications, provided the low-pressure permeate gas is used effectively. An added benefit compared with TEG units is the absence of BTEX emissions with membranes.
5.28 Other Processes
The first process is the Twister technology, which is discussed in Chapter 6. It has been considered attractive in offshore applications for dehydration because of its simplicity (no moving parts) along with its small size and weight.
Brouwer et al. (2004) discuss the successful implementation on an offshore platform. Some offshore field pressures are greater than 2,000 psi (140 bar), so recompression is not needed with the unit where overall pressure drop is 20 to 30%.
The second process is the vortex tube technology, which also is discussed in Chapter 6. It also has no moving parts. According to vendor information, it is used in Europe in conjunction with TEG addition to remove water from gas stored underground. We found no examples of its use in gas plants.
5.29 Comparison of Dehydration Processes
A number of factors should be considered in the evaluation of a dehydration process or combination of processes. If the gas must be dried for cryogenic liquids recovery, molecular sieve is the only long-term, proven technology available. It has the added advantage that it can remove CO2 at the same time. If CO2 is being simultaneously removed, because water displaces CO2, the bed must be switched before the CO2 breaks through, which is before any water breakthrough.
Enhanced TEG regeneration systems may begin to compete with molecular sieve. Skiff et al. (2003) claim to have obtained less than 0.1 ppmv water by use of TEG with a modified regeneration system that uses about 70% of the energy required for molecular sieves.
High inlet water-vapor concentrations make molecular sieve dehydration expensive because of the energy consumption in regeneration. Two approaches are used to reduce the amount of water going to the molecular sieve bed. First, another dehydration process, (e.g., glycol dehydration) is put in front of the molecular sieve bed. The second option is to have combined beds with silica gel or activated alumina in front of the molecular sieve. The bulk of the water is removed with the first adsorbent, and the molecular sieve removes the remaining water. This configuration reduces the overall energy required for regeneration.
If dehydration is required only to avoid free-water formation or hydrate formation or to meet the pipeline specification of 4 to 7 lb/MMscf (60 to 110 mg/Sm3), any of the above-mentioned processes may be viable. Traditionally, glycol dehydration has been the process of choice.
System constraints dictate which technology is the best to use. Smith (2004) provides an overview of natural gas dehydration technology, with an emphasis on glycol dehydration.
When considering susceptibility to inlet feed contamination, one should keep in mind that replacing a solvent is much easier and cheaper than changing out an adsorbent bed. However, prevention of contamination by use of properly designed inlet scrubbers and coalescing filters, if required, is the best solution.
In a conventional gas plant, where inlet fluctuations are handled in inlet receiving, feed contamination is generally limited to possible carryover from the sweetening unit. However, in field dehydration the possibility exists of produced water, solids, oil, and well-treating chemicals entering the dehydrator.
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5.22.11.4 Regeneration Gas Separator
Most desiccants also have an affinity for hydrocarbons, thus a skimmer is used to separate the valuable hydrocarbons from the water to be discarded. Frequent pH tests on the discarded water helps pinpoint corrosion problems in the adsorption system. A common problem encountered in regeneration gas separators is the fouling of the liquid dump line by desiccant dust and heavy oils.
5.22.11.5 Expander Plant Molecular Sieve Applications
Turbo expander plants commonly operate down to temperatures of -150 0F.
Operating points much below the equilibrium water content data illustrated in McKetta-Wehe chart (include designs to water contents as low as 1 ppm).
As shown in Table 5-7 (UP), only molecular sieves and activated alumina are capable of such performance. Molecular sieves are used in approximately 95% of the dehydration equipment for this type of plant (a 4A molecular sieve has twice the adsorptive capacity of activated alumina).
5.23 Desiccant Performance
5.23.1 General Conditions
Desiccants decline in adsorptive capacity at different rates under varying operating conditions.
Desiccant aging is a function of many factors, including:
Number of cycles experienced
Exposure to any harmful contaminants present in the inlet stream.
The most important variable affecting the decline rate of desiccant capacity is the chemical composition of the gas or liquid to be dried.
Capacity of a new desiccant will decline slowly during the first few months in service because of cyclic heating, cooling, and netting. Desiccant capacity usually stabilizes at about 55 to 70% of the initial capacity.
5.23.2 Moisture Analyzer
The moisture analyzer is used to optimize the drying cycle time, where the drying time is always shortened as the desiccant ages. Both inlet and outlet moisture analyzer probes should be used. A probe extending approximately 2 feet upward into the bed from the outlet end is recommended because it allows a dehydration capacity test to be run without the risk of a water breakthrough.
5.23.3 Effect of Contaminants in Inlet Feed Stream
Compressor oils, corrosion inhibitors, glycols, amines, and other high-boiling contaminants cause a decline in desiccant capacity, because normal reactivation temperatures will not vaporize the heavy materials. Residual contaminants slowly build up on the desiccant’s surface reducing the area available for adsorption. Many corrosion inhibitors chemically attack certain desiccants, permanently destroying their usefulness.
5.23.4 Effect of Regeneration Gases Rich in Heavy Hydrocarbons
Use of this rich gas in a 5500 to 600 0F regeneration service aggravates coking problems. Rich gases may be dried satisfactorily with molecular sieves. Lean dry gas is always preferable for regeneration.
5.23.5 Effect of Methanol in the Inlet Gas Stream
Methanol in the inlet gas is a major contributor to the coking of molecular sieves where regeneration is carried out at temperatures above 550 0F. Polymerization of methanol during regeneration produces dimethyl ether and other intermediates that will cause coking of the beds.
Conversion to ethylene glycol injection, instead of methanol for hydrate control, will increase sieve life and add at least 10% to sieve capacity by removing the vapor phase methanol from the system.
5.23.6 Useful Life
The most common reasons for replacing a bed are loss of adsorbent capacity and unacceptable pressure drop, which usually occur simultaneously. Values for the loss of capacity with time vary considerably, but common values used for molecular sieves in dehydration service are a 35% capacity loss over a 3 to 5 year period or a 50% loss in approximately 1,600 cycles. Typically, a rapid loss occurs in the beginning and a gradual loss thereafter. The adsorbent decays primarily because of carbon and sulfur fouling and caking caused by instability in the clay binder. These effects occur during bed regeneration.
Increased pressure drop is usually caused by breakdown of adsorbent into finer particles and by caking at the top of the bed because of refluxing. Attrition can occur when the pressure is increased or decreased (more than the allowable pressure-changing rate) after or before regeneration.
Monitoring the pressure drop is important, as it provides a good diagnostic to bed health.
In general, adsorbent life ranges from one to four years in normal service. Longer life is possible if feed gas is kept clean. Effectiveness of regeneration plays a major role in retarding the decline of a desiccant’s adsorptive capacity and prolonging its useful life. If all the water is not removed from the desiccant during each regeneration, its usefulness will sharply decrease.
5.23.7 Effect of Insufficient Reactivation
Insufficient reactivation can occur if the regeneration gas temperature or velocity is too low.
A desiccant manufacturer will generally recommend the optimum regeneration temperature and velocity for the product. Velocity should be high enough to remove the water and other contaminants quickly, thus minimizing the amount or residual water and protect the desiccant.
5.23.8 Effect of High Reactivation Temperature
Higher reactivation temperatures remove volatile contaminants before they form coke on the desiccant, maximizes desiccant capacity and ensures minimum effluent moisture content.
Final effluent hot gas temperature should be held one or two hours to achieve effective desiccant
reactivation.
5.24 Design
5.24.1 Pressure Drop & Minimum Diameter
The first step is to determine the bed diameter, which depends on the superficial velocity. Too large a diameter will require a high regeneration gas rate to prevent channeling. Too small a diameter will cause too high a pressure drop and damage the sieve. The pressure drop is determined by a modified Ergun equation, which relates pressure drop to superficial velocity as follows:
ΔP / L = B ս V + C ρ V2 Eq 5-17
ΔP = pressure drop, psi
L = length of packed bed, ft
B & C = constants from table 5-10
µ= viscosity, cp
V = superficial vapor velocity, ft/min
ρ = density, lb/ft3
Particle Type B C
1/8" bead (4x8 mesh) 0.0560 0.0000889
1/8" extrudate 0.0722 0.000124
1/16" bead (8x12 mesh) 0.152 0.000136
1/16" extrudate 0.238 0.000210
Table. 5-10. Constants for Equation 5-6.
Fig. 5-55 was derived from Eq 5-17 by assuming a gas composition and temperature and setting the maximum allowable ΔP/L equal to 0.33 psi/ft. The design pressure drop through the bed should be about 5 psi. A design pressure drop higher than 8 psi is not recommended as the desiccant is fragile and can be crushed by the total bed weight and pressure drop forces.
Fig. 5-55. Allowable Velocity for Mole Sieve Dehydrator
Remember to check the pressure drop after the bed height has been determined. Once the allowable superficial velocity is estimated, calculate the bed minimum diameter
(i.e., Dminimum), and select the nearest standard diameter (i.e. Dselected):
Dminimum = (4q /π Vmax)0.5 Eq. 5-18
where
D = diameter, ft
q = actual gas flow rate, ft3/min
q = m/60ρ Eq. 5-19
where
m = mass flow rate, lb/hr
ρ= density, lb/ft3
Obtain the corresponding superficial velocity, Vadjusted as follows:
Vadjusted = Vmax (Dminimum/ Dselected)2 Eq. 5-20
An alternative and more exact method to calculate the maximum superficial velocity can be determined by Eq 5-21, which was derived from Eq.5-17.
Vmax = [(ΔP/L)max/(Cρ)]0.5 – [(B/C)(ս/ρ)/2] Eq 5-21
The value of (ΔP/L)max in these equations depends on the sieve type, size, and shape, but a typical value for design is 0.33 psi/ft.
5.24.2 Mass desiccant Required & Bed Length
Choose an adsorption period and calculate the mass of desiccant required. Eight to twelve hour adsorption periods are common. Periods of greater than 12 hours may be justified especially if the feed gas is not water saturated.
Long adsorption periods mean less regenerations and longer sieve life, but larger beds and additional capital investment.
During the adsorption period, the bed can be thought of as operating with three zones.
Saturation or equilibrium zone
The middle or mass transfer zone (MTZ) is where the water content of the gas is reduced from its inlet concentration to < 1 ppm. (lb/MMscf ~ ppmv / 21.4)
The bottom zone is unused desiccant and is often called the active zone. If the bed operates too long in adsorption, the mass transfer zone begins to move out the bottom of the bed causing a “breakthrough.”
In the saturation zone, molecular sieve is expected to hold approximately 13 pounds of water per 100 pounds of sieve. New sieve will have an equilibrium capacity near 20%; 13% represents the approximate capacity of a 3-5 year old sieve.
This capacity needs to be adjusted when the gas is not water saturated or the temperature is above 75°F. Figures. 5-56 and 5.57 are used to find the correction factors for molecular sieve.
Alternatively, both parameters may be calculated using the following corellations:
CSS = 0.636 + 0.0826 ln(Sat) Eq. 5-22
and
CT = 1.20 – 0.0026 t(°F) Eq. 5-23
CT = 1.11 – 0.0047 t(°C) Eq. 5-24
where CSS and CT are correction factors for subsaturation and temperature, respectively.
Sat is the percent of saturation.
To determine the mass of desiccant required in the saturation zone, calculate the amount of water to be removed during the cycle and divide by the effective capacity.
SS = Wr / [(0.13)(CSS)(CT)] Eq 5-25
LS = 4 SS /[π (D2) (bulk density) Eq 5-26
Where
LS = length of packed bed saturation zone, ft
SS = amount molecular sieve required in saturation zone, lb
Wr = water removed per cycle, lb
CSS = saturation correction factor for sieve (Fig.5-56)
CT = temperature correction factor (Fig.5-57)
LS = length of packed bed saturation zone, ft
D = diameter, ft
Molecular sieve bulk density is 42 to 46 lb/ft3 for spherical particles and 40 to 44 lb/ft3 for extruded cylinders.
Even though the MTZ will contain some water (approximately 50% of the equilibrium capacity), the saturation zone is estimated assuming it will contain all the water to be removed.
The length of the mass transfer zone can be estimated as follows:
LMTZ = (Vadjusted/35)0.3 (ZL) Eq 5-27
Where:
LMTZ = length of packed bed mass transfer zone, ft
ZL = 1.70 ft for 1/8 inch sieve
= 0.85 ft for 1/16 inch sieve
Fig. 5-56. Mole Sieve Capacity Correction for gas percent saturation
Fig. 5-57. Mole Sieve Capacity Correction for Temperature
The total bed height is the summation of the saturation zone and the mass transfer zone heights. It should be no less than the vessel inside diameter, or 6 feet, whichever is greater.
Now the total bed pressure drop is checked. The ΔP/L for the selected diameter, Dselected, is adjusted using Eq 5-17 or the following approximation:
(ΔP / L)adjusted ≅ (0.33 psi/ft) (Vadjusted/Vmax)2 Eq 5-28
The result is multiplied times the total bed height (LS +LMTZ) to get the total design pressure drop, which should be 5- 8 psi. This is important, because the operating pressure drop can increase to as much as double the design value over three years. Too high a pressure drop plus the bed weight can crush the sieve. If the design pressure drop exceeds 8 psi, the bed diameter should be increased and the sieve amount and vessel dimensions recalculated.
To estimate the total cylindrical length of a tower, add 3 feet to the bed height, which provides the space for an inlet distributor and for bed support and hold-down balls under and on top of the sieve bed.
5.24.3 Regeneration Calculations
The first step is to calculate the total heat required to desorb the water and heat the desiccant and vessel. A 10% heat loss is assumed.
Qw = (1800 Btu/lb) (lbs of water on bed) Eq 5-29
Qsi = (lbs of sieve)(0.24 Btu/lb 0F) (Trg – Ti) Eq 5-30
Qst = (lbs of steel)(0.12 Btu/lb 0F) (Trg – Ti) Eq 5-31
Qhl =heat loss = (Qw +Qsi+Qst) (0.1) Eq 5-32
where
Qw = desorption of water heat duty, Btu
Qsi = duty required to heat mole sieve to regeneration temperature, Btu
Qst = duty required to heat vessel and piping to regeneration temperature, Btu
Qhl = regeneration heat loss duty, Btu
Trg = regeneration temperature, 0F
Ti = initial, temperature, 0F
The temperature, Trg, is the temperature to which the bed and vessel must be heated based on the vessel being externally insulated (i.e., no internal insulation which is usually the case). This is about 50°F below the temperature of the hot regeneration gas entering the tower.
The weight of the vessel steel is estimated from equations 5-33 and 5-34.
The design pressure, Pdesign, is usually set at 110% of the maximum operating pressure. The value of 0.125 in Eq 5-34 is the corrosion allowance in inches. The term 0.75Dbed is to account for the weight of the tower heads. The value of "3" provides the space for the inlet distributor and support and hold-down balls.
t(inches) = (12DbedPdesign) / (37,600 – 1.2Pdesign) Eq 5-33
Weight of steel (lb) = 155 (t + 0.125) (LS + LMTZ + 0.75Dbed + 3)Dbed Eq 5-34
For determination of the regeneration gas rate, calculate the total regeneration load from Eq. 5-35
Qtr = (2.5)(Qw + Qsi + Qst + Qhl) Eq. 5-35
where
Qtr = total regeneration heat duty, Btu
The 2.5 factor corrects for the change in temperature difference (in – out) across the bed with time during the regeneration period. It assumes that 40% of the heat in the gas transfers to the bed, vessel steel, and heat loss to atmosphere; and the balance leaves with the hot gas.
The regeneration-gas flow rate is calculated from Eq 5-36 below. The heat capacity, Cp, is calculated with Eq 5-37, with the enthalpies obtained from the enthalpy vs. temperature plots for various pressures in (GPSA- Section 24 – or refer to definition of (Cp) in eq. 5-37). The temperature, Thot, is 50°F above the temperature, Trg, to which the bed must be heated. The temperature, Tb, is the bed temperature at the beginning of regeneration, which is the same as the dehydration-plant feed temperature.
The heating time is usually 50% to 60% of the total regeneration time which must include a cooling period. Figure 5-59 shows a typical temperature profile for a regeneration period (heating and cooling). For 8 hour adsorption periods, the regeneration normally consists of 4 1/2 hours of heating, 3 hours of cooling and 1/2 hour for standby and switching. For longer periods the heating time can be lengthened as long as a minimum pressure drop of 0.01 psi/ft is maintained to ensure even flow distribution across the bed.
mrg = Qtr/[Cp(Thot -Tb)(heating time)] Eq 5-36
Cp (Btu/lb/°F) = (Hhot - Hi)/(Thot -Tb) Eq 5-37
Where
mrg = regeneration mass flow rate, lb/hr
Thot = Hot gas temperature, 0F
Tb = bed starting temperature, 0F
Cp = gas heat capacity, Btu/(lb . °F) = (Hhot - Hi)/(Thot -Tb) (Extract the value from fig. 5-60 ,which include values at 600 psia, or (Use Cp = 0.66 as an approximate value for natural gas in regeneration operation ranges) Hhot = enthalpy, BTU/lb, and Hi = enthalpy, BTU/lb. (Other values can be extracted from curves in GPSA, chapter 24 “Total Enthalpy of Paraffin Hydrocarbon Vapor”).
The superficial velocity of the regeneration gas is calculated from Eq 5-38 for which (q) is calculated from Eq 5-19
V = 4q / (π D2) Eq 5-38
where
V = Velocity, ft/sec
D = diameter, ft
q = actual gas flow rate, ft3/min
The calculated superficial velocity can not be less than the value that corresponds with a minimum bed pressure drop of 0.01 psi/ft. This can be determined from Fig. 5-58, which was derived from Eq 5-17 by assuming a gas composition and temperature and setting ΔP/L equal to 0.01 psi / ft.
If the calculated velocity is less than this, the regeneration gas rate, (mrg) must be increased by multiplying it by the ratio Vmin / V, and the period of regeneration should be decreased by multiplying it times the ratio V / Vmin. A more exact method for calculating the minimum superficial velocity is to use Eq 5-21, but to consider it in terms of (ΔP/L)min and Vmin instead of (ΔP/L)max and Vmax.
Fig. 5-58. Minimum Regeneration Velocity for Mole Sieve Dehydrator
Fig. 5-59. Inlet and Outlet Temperatures During Typical Solid Desiccant Bed Regeneration Period
5.24.4 Design Example
100 MMscfd of natural gas with a molecular weight of 18 is water saturated at 600 psia and 100°F and must be dried to –150°F dew point.
Determine the water content of the gas, and the amount of water that must be removed; and do a preliminary design of a molecular-sieve dehydration system consisting of two towers (one dehydration and the other regeneration in a cycle basis). Use 4A molecular sieve of 1/8" beads (i.e., 4x8 mesh). Compressibility factor z = 0.93.
The regeneration gas is part of the plant’s residue gas, which is at 600 psia and 100°F and has a molecular weight of 17. The bed must be heated to 500°F for regeneration. Compressibility factor = 0.95.
Solution Steps
1. Determine the bed diameter and the corresponding ΔP/L and V. First determine the maximum superficial velocity from Eq 5-10. Let the maximum ΔP/L be 0.33 psi/ft.
Vmax = [(ΔP/L)max/(Cρ)]0.5 – [(B/C)(ս/ρ)/2] Eq 5-21
Fro chapter 1, Eq. 1-9
ρg= 0.093 ((MW)P)/TZ lb/ft3 (Eq. 1-19)
ρg = 0.093 x 18 x 600 / (560 x 0.93)
ρ = 1.93 lb/ft3
Or,
ρ = density of gas (ρ)= (18 mole weight) (600 psia) / [10.73 (560 °R)(0.93 z)] = 1.93 lb/ft3
µ = 0.015 centipoise (Fig. 1-11)
C = 0.000089 from table 5-10.
Vmax = {(0.33 psi/ft) / (0.0000889)(1.93) lb/ft3)}0.5 – [(0.056 / 0.0000889) (0.015 centipoise/1.93 lb/ft3)/2]
= 41.4 ft/min
Mass flow rate (m) = (100 X 106 scf/day) (18 lb/lb mole) / [(24 hr/day)(379.5 scf/lb mole)]
m = 198,000 lb/hr of wet gas
From Eq. 5-19
q = (198,000 lb/hr)/((60 min/hr)(1.93 lb/ft3)) = 1710 ft3/min of wet gas
From eq. 5-18
Dminimum = [(4(1710 ft3/min))/(3.14 X 41.4 ft/min)]0.5 = 7.25 ft
Round off upward to 7.5 ft diameter, for which V and ΔP/L are adjusted as follows (eq. 5-20):
Vadjusted = (41.4)(7.25/7.5)2 = 38.7 ft/min
(ΔP/L)adjusted = 0.33(38.7/41.4)2 = 0.29 psi/ft
2. Estimate the amount of water to be removed from the feed per cycle for each bed.
Base this on a 24-hour cycle consisting of 12 hours adsorbing and 12 hours regenerating (heating, cooling, standby, and valve switching).
From chapter 4, fig. 4-8, the water content at 600 psia and 100°F is 88 pounds/MMscf. The water content at a dew point of –150°F is essentially zero, so the water removed is the following:
w = (88-0 lb/MMscf)(100 MMscf/day) / (24 hr/day)
= 367 lb/hr of water removed
Wr = (367 lb/hr) (12 hr) = 4404 lb water removed per 12-hour drying period or 24-hour cycle per bed.
3. Determine the amount of sieve required and the bed height based on a sieve bulk density of 45 lb/ft3 (table 5-7). Since the feed gas is water saturated, the relative humidity is 100%, so CSS is 1.0, and CT is 0.93 at 100°F from Figures. 5-56 & 5-57.
Applying the equations 5-25:
SS = Wr / [(0.13)(CSS)(CT)] Eq 5-25
SS = (4404) / ((0.13)(1.0)(0.93)) = 36,427 lb of sieve for each bed.
From eq. 5-26.
LS = 4 SS /[π (D2) (bulk density) Eq 5-26
LS = (4)(36,427) /((3.1416) (7.5)2 (45)) = 18.3 ft bed height
from eq-5-27
LMTZ = (Vadjusted/35)0.3 (ZL) Eq 5-27
LMTZ = (38.7/35)0.3 (1.7) = 1.8 ft for mass-transfer zone
LS + LMTZ = 18.3 + 1.8 = 20.1 ft of sieve for each bed
The total sieve = (20.1/18.3)(36,427) = 40,010 lb for each bed
4. Check the bed design and pressure drop which is the ΔP/L calculated in Step 1 times the total bed height calculated in Step 3:
(0.29 psi/ft)(20.1 ft) = 5.8 psi which meets the criterion of not exceeding 8 psi.
5. Calculate the total heat required to desorb the water based on heating the bed and vessel to 500°F. First calculate the weight of steel from Eq 5-33 and 5-34. Let the design pressure, Pdesign, be 110% of the operating pressure: Pdesign = (600)(1.1) = 660 psia.
t(inches) = (12DbedPdesign) / (37,600 – 1.2Pdesign) Eq 5-33
Weight of steel (lb) = 155 (t + 0.125) (LS + LMTZ + 0.75Dbed + 3)Dbed Eq 5-34
t = (12)(7.5)(660) / (37,600 – (1.2)(660)) = 1.614 inches
Weight of steel = (155) (1.614 + 0.125) (18.3 + 1.
+ (0.75) (7.5) + 3) (7.5) = 58,070 poundsFrom the following equations:
Qw = (1800 Btu/lb) (lbs of water on bed) Eq 5-29
Qsi = (lbs of sieve)(0.24 Btu/lb 0F) (Trg – Ti) Eq 5-30
Qst = (lbs of steel)(0.12 Btu/lb 0F) (Trg – Ti) Eq 5-31
Qhl =heat loss = (Qw +Qsi+Qst) (0.1) Eq 5-32
Qtr = (2.5)(Qw + Qsi + Qst + Qhl) Eq. 5-35
Qw = (1800 Btu/lb (4404 lb water) = 7,927,000 Btu
Qsi = (40,010 lb) (0.24 Btu/lb/°F) (500°F – 100°F) = 3,841,000 Btu
Qst = (58,070 lb) (0.12 Btu/lb/°F) (500°F – 100°F) = 2,787,000 Btu
Qhl = (2,787,000 + 7,927,000 + 3,841,000) (0.10) = 1,455,000 Btu
Qtr = (2.5) (2,787,000 + 7,927,000 + 3,841,000 + 1,455,000) = 40,025,000 Btu
6. Calculate the flow rate of regeneration gas using Eq 5-36. Let the heating time be 60% of the total regeneration period, and calculate the gas heat capacity, use Cp, = 0.66 (Btu/lb/°F), or (calculate it from Eq 5-37 using enthalpy curves in GPSA, chapter 24 “Total Enthalpy of Paraffin Hydrocarbon Vapor”):
mrg = Qtr/[Cp(Thot -Tb)(heating time)] Eq 5-36
Cp (Btu/lb/°F) = (Hhot - Hi)/(Thot -Tb) Eq 5-37
(60%) (12 hr) = 7.2 hours heating
Cp (@ 600 psia from fig 5-60) = ((545 – 250) (Btu/lb))/((550 – 100) (°F)) = 0.66 Btu/lb/°F
mrg = (40,025,000 Btu)/((0.66 Btu/lb/°F) (550 – 100) (°F) (7.2 hr)) = 18,717 lb/hr (Eq 5-22)
7. Check that the ΔP/L ≥ 0.01 psi/ft at 550°F.
ρg = 0.093 x 17 x 600 / (1110 x 0.95)
ρ = 0.9 lb/ft3
From Eq. 5-19 q = m/60ρ
= 18,717 / (60 x 0.9) = 346.6 ft3/min of hot regeneration gas
Rearranging Eq 5-38:
V = 4q / (π D2) Eq 5-38
V = 4q/πD2 = ((4)(346.6)/((3.414)(7.5)2) = 7.21 ft/min
μ = 0.023 cP (Fig. 1-11)
From ΔP / L = B ս V + C ρ V2 Eq 5-17
ΔP/L = (0.056) (0.023) (7.21) + (0.0000889) (0.9) (7.21)2 = 0.013 psi/ft (Eq 5-6)
This is safely above the minimum value of 0.01 psi/ft needed to prevent channeling.
8. The design results are summarized as follows:
Number of vessels: two
Vessel design pressure and temperature: 660 psig and 600°F
Vessel dimensions: 90 inches (7.5 feet) ID by 23.1 feet tan to tan
Weight of molecular sieve: 2x40,010 lb
Regeneration gas rate: 18,717 lb/hr (10.026 MMscfd)
Regeneration gas temperature: 550°F
Cycle time: 24 hours, 12 hours adsorption, 12 hours regeneration
[img]ttp://oilprocessing.net/data/documents/V5-60.png[/img]
Fig.5-60. Total Enthalpy of Paraffin Hydrocarbon Vapor @ 600 psia.
5.25 Nonregenerable Dehydrator
In some situations, such as remote gas wells, use of a consumable salt desiccant, such as CaCl2, may be economically feasible.
5.25.1 Calcium Chloride Dehydrator Unit
Calcium chloride (CaCI2) dehydrator consists of three sections (Figure 5-61):
Inlet gas scrubber
Brine tray
Solid brine particles
5.25.2 Principles of Operation
Solid desiccant is placed in the top of the unit. Water-wet gas contacting the solid CaCI2 gives up part of its water to form liquid brine to drip down and fill the trays.
Solid anhydrous CaCl2 combines with water to form various CaCl2 hydrates (CaCl2 .XH2O). As water absorption continues, CaCl2 is converted to successively higher states of hydration eventually forming a CaCl2 brine solution.
Inlet gas coming up through the specially designed nozzles on the trays contact the brine efficiently. The wettest gas contacts the most dilute brine (about 1.2 specific gravity).
Approximate 2.5 Lb H2O/lb CaCI2 is removed in the trays. Brine gravity on the top tray is about 1.4. Another 1 lb H2O/lb CaCI2 is removed in the solid bed section.
Outlet water contents of 1 lb/MMscf have been achieved with CaCl2 dehydrators. Typical CaCl2
Maximum dew point depression of 60 0 to 70 0F occurs in this section. Typically used in remote, small gas fields without heat or fuel. capacity is 0.3 lb CaCl2 per lb H20. Superficial bed velocities are 20-30 ft/min and length to diameter ratio for the bed should be at least 3 to 4:1.
CaCl2 dehydrators may offer a viable alternative to glycol units on low rate, remote dry gas wells. The CaCl2 must be changed out periodically. In low capacity – high rate units this may be as often as every 2-3 weeks. Brine disposal raises environmental issues. In addition, under certain conditions the CaCl2 pellets can bond together to form a solid bridge in the fixed bed portion of the tower. This results in gas channeling and poor unit performance.
Advantages Disadvantages
Simple
No moving parts
No heat required
Does not react with H2S or CO2
Can dehydrate hydrocarbon liquids Batch process
Emulsifies with oil
Unreliable
Limited dew point depression
Table. 5-11. Advantages and disadvantage of Calcium Chloride Dehydrator Unit
Operating Problems
Bridging and channeling is a problem.
Brine can crystallize at 85 0F, thus during low flow periods can plug vessel outlet or trays.
Brine carry-over can cause severe corrosion problems.
Fig. 5-61. CaCl2 dehydrator.
Design Considerations
Figure 5-62 illustrates the water content of natural gas dried by solid calcium chloride bed units.
Fig. 5-62 Water content on natural gas driven by CaCl2 unit (Left: freshly recharged; right: just prior to recharging).
5.26 Dehydration by Refrigeration
The dehydration of natural gas can also be achieved by refrigeration and/or cryogenic processing down to – 150°F in the presence of methanol hydrate and freeze protection. The condensed water and methanol streams decanted in the cold process can be regenerated by conventional distillation or by a patented process called IFPEX-1®.
In the latter process illustrated in schematic form in Figure 5-63 a slip stream of water saturated feed gas strips essentially all the methanol in the cold decanted methanol water stream originating in the cold process at feed gas conditions to recirculate the methanol to the cold process. The water stream leaving the stripper contains generally less than 100 ppm wt of methanol. No heat is required for the process and no atmospheric venting takes place.
The process has several major advantages:
• It can obtain dew points in the −100 to −150°F (−70 to –100°C) range.
• It requires no heat input other than to the methanol regenerator.
• It requires no venting of hydrocarbon-containing vapors.
However, it requires external refrigeration to cool the gas, and minimal methanol losses occur in the stripper.
Fig. 5-63. Example IFPEX-1 . Dehydration Process Flow Diagram
5.27 Dehydration by Membrane Permeation
Membranes can be used to separate gas stream components in natural gas such as water, CO2 and hydrocarbons according to their permeabilities. Each gas component entering the separator has a characteristic permeation rate that is a function of its ability to dissolve in and diffuse through the membrane.
The driving force for separation of a gas component in a mixture is the difference between its partial pressure across the membrane. As pressurized feed gas flows into the metal shell of the separator, the fast gas component, such as water and CO2, permeate through the membrane. This permeate is collected at a reduced pressure, while the non-permeate stream, i.e., the dry natural gas, leaves the separator at a slightly lower pressure than the feed.
The amount of methane and other natural gas components in the permeate stream is dependent on pressure drop and the surface area of the membranes. However, 5–10% of the feed stream is a realistic figure. Dehydration by membrane permeation is therefore normally only considered for plants that can make use of low pressure natural gas fuel.
Membranes are characteristic by lightweight, large turndown ratio, and low maintenance, that make them competitive with glycol units in some situations.
Feed pretreatment is a critical component of a membrane process. The inlet gas must be free of solids and droplets larger than 3 microns.
Inlet gas temperature should be at least 20°F (10°C) above the dew point of water to avoid condensation in the membrane.
Units operate at pressures up to 700 to 1,000 psig (50 – 70 barg) with feed gases containing 500 to 2,000 ppmv of water (lb/MMscf ~ ppmv / 21.4). They produce a product gas stream of 20 to 100 ppmv and 700 to 990 psig (48 to 68 barg). The low-pressure (7 to 60 psig [0.5 to 4 barg]) permeate gas volume is about 3 to 5% of the feed gas volume.
This gas must be recompressed or used in a low-pressure system such as fuel gas.
Smith (2004) suggests that membranes used for natural gas dehydration are economically viable only when dehydration is combined with acid-gas removal.
On the basis of commercial units installed and several studies; (Bikin et al., 2003), membranes are economically attractive for dehydration of gas when flow rates are less than 10 MMscfd (0.3 MMSm3/d). Binci et al. claim that membrane units are competitive with TEG dehydrators on offshore platforms at flows below 56 MMscfd (1.6 MMSm3/d). Certainly, the reliability and simplicity of membranes make them attractive for offshore and remote-site applications, provided the low-pressure permeate gas is used effectively. An added benefit compared with TEG units is the absence of BTEX emissions with membranes.
5.28 Other Processes
The first process is the Twister technology, which is discussed in Chapter 6. It has been considered attractive in offshore applications for dehydration because of its simplicity (no moving parts) along with its small size and weight.
Brouwer et al. (2004) discuss the successful implementation on an offshore platform. Some offshore field pressures are greater than 2,000 psi (140 bar), so recompression is not needed with the unit where overall pressure drop is 20 to 30%.
The second process is the vortex tube technology, which also is discussed in Chapter 6. It also has no moving parts. According to vendor information, it is used in Europe in conjunction with TEG addition to remove water from gas stored underground. We found no examples of its use in gas plants.
5.29 Comparison of Dehydration Processes
A number of factors should be considered in the evaluation of a dehydration process or combination of processes. If the gas must be dried for cryogenic liquids recovery, molecular sieve is the only long-term, proven technology available. It has the added advantage that it can remove CO2 at the same time. If CO2 is being simultaneously removed, because water displaces CO2, the bed must be switched before the CO2 breaks through, which is before any water breakthrough.
Enhanced TEG regeneration systems may begin to compete with molecular sieve. Skiff et al. (2003) claim to have obtained less than 0.1 ppmv water by use of TEG with a modified regeneration system that uses about 70% of the energy required for molecular sieves.
High inlet water-vapor concentrations make molecular sieve dehydration expensive because of the energy consumption in regeneration. Two approaches are used to reduce the amount of water going to the molecular sieve bed. First, another dehydration process, (e.g., glycol dehydration) is put in front of the molecular sieve bed. The second option is to have combined beds with silica gel or activated alumina in front of the molecular sieve. The bulk of the water is removed with the first adsorbent, and the molecular sieve removes the remaining water. This configuration reduces the overall energy required for regeneration.
If dehydration is required only to avoid free-water formation or hydrate formation or to meet the pipeline specification of 4 to 7 lb/MMscf (60 to 110 mg/Sm3), any of the above-mentioned processes may be viable. Traditionally, glycol dehydration has been the process of choice.
System constraints dictate which technology is the best to use. Smith (2004) provides an overview of natural gas dehydration technology, with an emphasis on glycol dehydration.
When considering susceptibility to inlet feed contamination, one should keep in mind that replacing a solvent is much easier and cheaper than changing out an adsorbent bed. However, prevention of contamination by use of properly designed inlet scrubbers and coalescing filters, if required, is the best solution.
In a conventional gas plant, where inlet fluctuations are handled in inlet receiving, feed contamination is generally limited to possible carryover from the sweetening unit. However, in field dehydration the possibility exists of produced water, solids, oil, and well-treating chemicals entering the dehydrator.
Fundamentals of Oil and Gas Processing
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Basics of Gas Field Processing
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Re: Basics of Gas Field Processing Book "Full text"
Gas Sweetening and Sulfur Recovery - Chapter 6 - Part 1 A.
Chapter 6 216
Gas Sweetening 216
6.1 Introduction 216
6.1.1 Definition 216
6.1.2 Safety Precautions 216
6.1.3 Purification Levels 217
6.1.4 Acid Gas Disposal 217
6.1.5 Removal of contaminants and pretreatment 218
6.2 Gas Treating – Process Options 219
6.3 Chemical Solvent Processes 221
6.3.1 Amine (Aqueous Alkanolamine) Processes 223
6.3.2 Alkaline Salt Process (Hot Carbonate) 234
6.3.3 Specialty Batch Chemical Solvents 237
6.4 Physical Solvent Processes 238
6.4.1 Selexol® 239
6.4.2 Fluor Solvent 242
6.4.3 Rectisol Process® 243
6.4.5 Purisol® 243
6.4.5 Catasol® 243
6.4.6 Morphysorb® 243
6.5 Hybrid Solvent Processes 244
6.5.1 Sulfinol® Process 244
6.5.2 Selefining Process 244
6.6 General Considerations for Solvent Process 245
6.6.1 Solution Filtration 245
6.6.2 Flash Tank 245
6.6.3 Corrosion 246
6.6.4 Foaming 247
6.6.5 Materials 248
6.7 Solid Bed Processes 248
6.7.1 General Process Description 248
6.7.2 Iron Sponge Process 248
6.7.3 SulfaTreat® 250
6.7.4 Zinc Oxide Process 251
6.7.5 Chemsweet® 251
6.7.6 PuraSpec® 251
6.7.8 Molecular Sieve Process 251
6.7.9 Oxorbon 253
6.8 Direct Conversion Processes (Liquid Redox) 254
6.8.1 Stretford Process 255
6.8.2 Lo-Cat process 255
6.8.3 Sulferox process 257
6.8.4 IFP Process 258
6.9 Distillation Process 258
6.10 Sulfur Recovery (The Claus Process) 259
6.10.1 Claus Process Considerations 261
6.10.2 Process Variations 262
6.10.3 Combustion Operation 263
6.10.4 Claus Unit Tail Gas Handling 263
6.11 Gas Permeation Process (Membranes) 266
6.11.1 Membrane Fundamentals 266
6.11.2 Membrane Selection Parameters 266
6.11.3 Membrane Structure Types 267
6.11.4 Carbon Dioxide Removal from Natural Gas 268
6.11.5 Membrane Elements 268
6.11.6 Membrane Design Considerations 271
6.11.7 Operating Considerations 274
6.11.8 Feed Gas Pretreatment 277
6.11.9 Membrane Advantages & Disadvantages 280
6.11.10 Hybrid Configurations 282
6.12 Biological Processes 285
6.13 Process Selection 285
6.13.1 Inlet Gas Stream Analysis 285
6.13.2 General Considerations 285
6.13.3 Removal of H2S 286
6.13.4 Removal of H2S and CO2 286
6.13.5 Process Selection charts 287
6.14 Safety & Environmental Considerations 289
6.15 Design Procedure 290
6.15.1 Iron Sponge 290
Example 6-1 291
6.15.2 The Amine System 292
Method 1 292
Example 6-2 294
Method 2 294
Example 6-3 296
-----------------
Chapter 6
Gas Sweetening
6.1 Introduction
Gas treating involves reduction of the “acid gases” carbon dioxide (CO2) and hydrogen sulfide (H2S), along with other sulfur species, to sufficiently low levels to meet contractual specifications or permit additional processing in the plant without corrosion and plugging problems. This chapter focuses on acid gases because they are the most prevalent. When applicable, discussion is given to other sulfur species.
There are many methods that may be employed to remove acidic components (primarily H2S and CO2) and other impurities from hydrocarbon streams. The available methods may be broadly categorized as those depending on chemical reaction, absorption, adsorption or permeation. Processes employing each of these techniques will be described.
6.1.1 Definition
Acid Gases
H2S combined with water forms sulfuric acid. CO2 combined with water forms carbonic acid.
Both are undesirable because they cause corrosion and reduce heating value and sales value. H2S is poisonous and may be lethal.
Sour Gas:
Sour gas is defined as natural gas with H2S and other sulfur compounds.
Sweet Gas:
Sweet gas is defined as natural gas without H2S and other sulfur compounds.
Acid gases Partial Pressure:
Partial pressure is used as an indicator if treatment is required. Partial pressure is defined as
PP = (total pressure of system) x ( mol% of gas)
where CO2 is present with water, a partial pressure >30 psia, would indicate CO2 corrosion might be expected. Below 15 psia, would indicate CO2 corrosion would not normally be a problem although inhibition may be required.
Factors that influence CO2 corrosion are those directly related to solubility, that is, temperature, pressure, and composition of the water. Increased pressure increases solubility and increased temperature decreases solubility.
H2S may cause sulfide stress cracking due to hydrogen embrittlement in certain metals. H2S partial pressure >0.05 psia, necessitates treating.
6.1.2 Safety Precautions
Hydrogen sulfide is a highly toxic gas. At concentrations as low as 10 ppmv irritation of the eyes, nose, and throat is possible. The human nose can detect hydrogen sulfide in concentrations as low as 0.02 ppmv. However, the human sense of smell cannot be relied on to detect hazardous concentrations of hydrogen sulfide. Higher concentrations and extended exposure to hydrogen sulfide will desensitize the sense of smell. The concentrations required for different reactions by the human body are:
1. Threshold limit value (TLV) for prolonged exposure: 10 ppmv
2. Slight symptoms after several hours exposure: 10-100 ppmv
3. Maximum concentration that can be inhaled for one hour without serious effects such as significant eye and respiratory irritation: 200-300 ppmv
4. Dangerous after exposure of 30 minutes to one hour: 500-700 ppmv
5. Fatal in less than 30 minutes: 700-900 ppmv and above.
6. Death in minutes: greater than 1000 ppmv
Hydrogen sulfide is highly flammable and will combust in air at concentrations from 4.3 to 46.0 volume percent. Hydrogen sulfide vapors are heavier than air and may migrate considerable distances to a source of ignition.
When H2S concentrations are well above the ppmv level, other sulfur species can be present. These compounds include carbon disulfide (CS2), mercaptans (RSH), and sulfides (RSR), in addition to elemental sulfur. If CO2 is present as well, the gas may contain trace amounts of carbonyl sulfide (COS). The major source of COS typically is formation during regeneration of molecular-sieve beds used in dehydration (see Chapter 5).
Gaseous carbon dioxide is a naturally occurring gas that is 50% heavier than air and is colorless and odorless. It is also a principal by-product of combustion. CO2 is inactive and therefore non-flammable. CO2 will displace oxygen and can create an oxygen-deficient atmosphere resulting in suffocation. The principal hazard of CO2 is exposure to elevated concentrations.
The atmospheric concentration immediately hazardous to life is 10%. Because CO2 is heavier than air, its hazard potential is increased, especially when entering tanks and vessels.
"A common but erroneous belief is that CO2 simply acts as an asphyxiant by lowering the oxygen level below the 16 percent minimum necessary to sustain life (at sea level). Although this is frequently the case in most serious accidents, CO2 begins to have a noticeable effect on normal body functions at about two to three percent. The concentration of carbon dioxide in the blood affects the rate of breathing, a measurable increase resulting from a level of one percent in the inspired air."
6.1.3 Purification Levels
Although many natural gases are free of objectionable amounts of H2S and CO2, substantial quantities of these impurities are found in both gas reserves and production. In a survey of U.S. gas resources, Meyer (2000) defined subquality gas as that containing CO2 ≥ 2%, N2 ≥ 4%, or H2S ≥ 4 ppmv.
The inlet conditions at a gas processing plant are generally temperatures near ambient and pressures in the range of 300 to 1,000 psi (20 to 70 bar), so the partial pressures of the entering acid gases can be quite high. If the gas is to be purified to a level suitable for transportation in a pipeline and used as a residential or industrial fuel, then the H2S concentration must be reduced to 0.25 gr/100 scf (4 ppm, or 6 mg/m3), and the CO2 concentration must be reduced to a maximum of 3 to 4 mol%. However, if the gas is to be processed for NGL recovery or nitrogen rejection in a cryogenic turbo expander process, CO2 may have to be removed to prevent formation of solids. If the gas is being fed to an LNG liquefaction facility, then the maximum CO2 level is about 50 ppmv because of potential solids formation.
6.1.4 Acid Gas Disposal
What becomes of the CO2 and H2S after their separation from the natural gas?
The answer depends to a large extent on the quantity of the acid gases. For CO2, if the quantities are large, it is sometimes used as an injection fluid in EOR (enhanced oil recovery) projects. Several gas plants exist to support CO2 flooding projects. If this option is unavailable, then the gas can be vented, provided it satisfies environmental regulations for impurities.
In the case of H2S, four disposal options are available:
1. Incineration and venting, if environmental regulations regarding sulfur dioxide emissions can be satisfied
2. Reaction with H2S scavengers, such as iron sponge
3. Conversion to elemental sulfur by use of the Claus or similar process
4. Disposal by injection into a suitable underground formation.
The first two options are applicable to trace levels of H2S in the gas, and the last two are required if concentrations are too high to make the first two options feasible.
6.1.5 Removal of contaminants and pretreatment
All gas sweetening units should have well-designed pretreatment facilities. Carryover of brine or liquid hydrocarbon (as slugs or aerosol) from upstream production operations can cause problems for gas treating and downstream processing equipment. Also, field facilities are not typically designed to remove troublesome contaminants like gas-phase heavy hydrocarbons.
These contaminants can likewise cause operational difficulties in the sweetening process.
If gross liquid carryover from an upstream facility is possible, a slug catcher is recommended. It should be sized not only for steady inlet fluid volumes, but for surge capacity to handle slugs of liquid hydrocarbons, water, and/ or well treatment chemicals.
If aerosols are a concern, an inlet filter separator is suggested. Selected filter elements can remove entrained droplets down to 0.3 microns in diameter.
For liquid hydrocarbon treatment, a filter coalescer may be used to remove suspended water or glycol prior to further processing.
Heavy hydrocarbons (C6+) can be absorbed by solvents, which could lead to foaming in the sweetening unit. It is possible to reduce the heavy hydrocarbon content of the incoming gas through cooling (via Joule-Thomson expansion, propane refrigeration, or turbo-expansion), and subsequent condensation of the heavy components. The condensed liquids are removed, and the gas is warmed above the saturation temperature before going to the sweetening unit.
An alternative means to remove gaseous heavy hydrocarbons is through adsorption. Either alumina or silica gel beds may be used in parallel such that one bed is regenerated while the other is in service. The beds are regenerated by heating and desorbing the hydrocarbons. The heavy hydrocarbons are recovered from the regeneration gas via condensation.
Oxygen entry into a hydrocarbon system is often troublesome. If liquid water is present, severe corrosion may occur. If H2S or sulfur is present, corrosion by a different mechanism or sulfur deposition and plugging may occur. Oxygen contamination may be addressed by several different approaches but the first step is to find and correct the source of oxygen entry into the system. This is often the simplest and most cost effective approach. Most oxygen leaks may be traced to compressor suctions or pipe fittings.
To eliminate oxygen contamination a number of possibilities exist:
• React the oxygen with chemicals.
Chemicals such as amines, organics or inorganic compounds may be added to remove free oxygen. Oxygen scavengers are available from many chemical suppliers.
• Thermal oxidation reactions.
Integrated processes such as the DeOxy by “Optimized Process Designs” can perform a very limited burn to consume the free oxygen.
• Remove other reactants that cause problems with the presence of oxygen.
By removing offending components such as water or H2S that react with oxygen, the presence of low amounts of oxygen may be tolerated.
• Treat the symptom.
Corrosion inhibitors, filtration and / or alternate schemes may be utilized to stop or offset any adverse effects of oxygen contamination.
6.2 Gas Treating – Process Options
Sulfur exists in natural gas as hydrogen sulfide (H2S), and the gas is usually considered sour if the hydrogen sulfide content exceeds 4 ppm of H2S. The process for removing hydrogen sulfide and carbon dioxide from a natural gas stream is referred to as “sweetening” the gas.
Numerous processes have been developed for acid gas removal and gas sweetening based on a variety of chemical and physical principles.
Some of the more important items that must be considered before a process is selected are:
• The type and concentration of impurities and hydrocarbon composition of the sour gas. For example, COS, CS2, and mercaptans can affect the design of both gas and liquid treating facilities. Physical solvents tend to dissolve heavier hydrocarbons, and the presence of these heavier compounds in significant quantities tends to favor the selection of a chemical solvent.
• The temperature and pressure at which the sour gas is available. High partial pressures (50 psi [3.4 bar] or higher) of the acid gases in the feed favor physical solvents, whereas low partial pressures favor the amines.
• The specifications of the outlet gas (low outlet specifications favor the amines).
• The volume of gas to be processed.
• The specifications for the residue gas, the acid gas, and liquid products.
• The selectivity required for the acid gas removal.
• The capital, operating, and royalty costs for the process.
• The environmental constraints, including air pollution regulations and disposal of byproducts considered hazardous chemicals.
If gas sweetening is required offshore, both size and weight are additional factors that must be considered. Whereas CO2 removal is performed offshore, H2S removal is rarely done unless absolutely necessary because of the problems of handling the rich acid gas stream or elemental sulfur.
If the gas processing facility is to be used in conjunction with liquids recovery, the requirements for H2S, CO2, and mercaptan removal may be affected. In liquid recovery plants, varying amounts of H2S, CO2, and other sulfur compounds will end up in the liquid product. Failure to remove these components prior to liquids recovery may require liquid product treating in order to meet product specifications. In many instances, liquid treating may be required anyway.
When sulfur recovery is required, the composition of the acid gas stream feeding the sulfur plant must be considered. With CO2 concentrations greater than 80% in the acid gas, the possibility of selective treating should be considered to raise the H2S concentration to the sulfur recovery unit (SRU). This may involve a multi-stage gas treating system in which the gas exiting the first stage is enriched by passing it through another absorption solvent loop.
High concentrations of hydrocarbons can cause design and operating problems for the SRU. The effect of these components must be weighed when selecting the gas treating process to be used.
Decisions in selecting a gas treating process can be simplified by gas composition and operating conditions.
High partial pressures (50 psi) of acid gases enhance the possibility of using a physical solvent. The presence of significant quantities of heavy hydrocarbons in the feed discourages using physical solvents. Low partial pressures of acid gases and low outlet specifications generally require the use of amines for adequate treating. Process selection is not easy and a number of variables must be weighed prior to making a process selection.
Figure 6-1 summarizes the more important processes and groups them into the generally accepted categories.
Fig. 6-1. Acid Gas Removal Processes
Table 6.1 lists the processes used to separate the acid gas from other natural gas components.
Tables 6-2 and 6-3, include comparison of different gas sweetening processes.
Table 6.1 Acid Gas Removal Processes
Note a: MEA reacts nonreversibly with COS (carbonyl sulfide), and, therefore, should not be used to treat gases with a large concentration of COS.
Table 6.2 Gases Removed by Various Processes
6.3 Chemical Solvent Processes
Chemical reaction processes remove the H2S and/or CO2 from the gas stream by chemical reaction with a material in the solvent solution. The reactions may be reversible or irreversible.
In reversible reactions the reactive material removes CO2 and/or H2S in the contactor at high partial pressure and/or low temperature. The reaction is reversed by high temperature and/or low pressure in the stripper. In irreversible processes the chemical reaction is not reversed and removal of the H2S and/or CO2 requires continuous makeup of the reacting material.
Fig. 6-6 shows the process flow for a typical reversible chemical reaction process.
In solvent absorption, the two major cost factors are the solvent circulation rate, which affects both equipment size and operating costs, and the energy require and disadvantages of chemical and physical solvents. Table 6.4 summarizes some of the advantages and disadvantages of chemical and physical solvents.
Table 6-3. CO2 and H2S Removal Processes for Gas Streams.
The most common chemical solvents are:
• Amines (Aqueous Alkanolamine)
• Carbonates
Table 6.4. Comparison of Chemical and Physical Solvents
6.3.1 Amine (Aqueous Alkanolamine) Processes
All commonly used amines are alkanolamines, which are amines with OH groups attached to the hydrocarbon groups to reduce their volatility. Figure 6.2 shows the formulas for the common amines used in gas processing.
Amines are compounds formed from ammonia (NH3) by replacing one or more of the hydrogen atoms with another hydrocarbon group. Replacement of a single hydrogen produces a primary amine, replacement of two hydrogen atoms produces a secondary amine, and replacement of all three of the hydrogen atoms produces a tertiary amine. Primary amines form stronger bases than secondary amines, which form stronger bases than tertiary amines. Amines with stronger base properties are more reactive toward CO2 and H2S gases and form stronger chemical bonds. Sterically hindered amines are compounds in which the reactive center (the nitrogen) is partially shielded by neighboring groups so that larger molecules cannot easily approach and react with the nitrogen. The amines are used in water solutions in concentrations ranging from approximately 10 to 65 wt% amines.(table. 6-5 lists Amines physical properties ).
Table. 6-6. lists approximate guidelines for a number of alkanolamine processes.
The amine processes are particularly applicable where acid gas partial pressures are low and/or low levels of acid gas are desired in the residue gas.
Because the water content of the solution minimizes heavy hydrocarbon absorption, these processes are well suited for gases rich in heavier hydrocarbons. Some amines can be used to selectively remove H2S in the presence of CO2.
Amines remove H2S and CO2 in a two step process:
1. The gas dissolves in the liquid (physical absorption).
2. The dissolved gas, which is a weak acid, reacts with the weakly basic amines.
Absorption from the gas phase is governed by the partial pressure of the H2S and CO2 in the gas, whereas the reactions in the liquid phase are controlled by the reactivity of the dissolved species. The principal reactions are summarized in the next section.
Fig. 6.2 Molecular structures of commonly used amines.
Table 6-5 physical properties of gas sweetening chemicals.
Table 6-6. Approximate Guidelines for Amine Processes
Table. 6-5. Physical properties of gas treating chemicals.
Table 6-6. Approximate Guidelines for Amine Processes
Notes:
1. These data alone should not be used for specific design purposes. Many design factors must be considered for actual plant design.
2. Dependent upon acid gas partial pressures and solution concentrations.
3. Dependent upon acid gas partial pressures and corrosiveness of solution. Might be only 60% or less of value shown for corrosive systems.
4. Varies with stripper overhead reflux ratio. Low residual acid gas contents require more stripper trays and/or higher reflux ratios yielding larger reboiler duties.
5. Varies with stripper overhead reflux ratios, rich solution feed temperature to stripper and reboiler temperature.
6. Maximum point heat flux can reach 20,000–25,000 Btu/hr-ft2 at highest flame temperature at the inlet of a direct fired fire tube. The most satisfactory design of firetube heating elements employs a zone by zone calculation based on thermal efficiency desired and limiting the maximum tube wall temperature as required by the solution to prevent thermal degradation. The average heat flux, Q/A, is a result of these calculations.
7. Reclaimers are not used in DEA and MDEA systems.
8. Reboiler temperatures are dependent on solution conc. flare/vent line back pressure and/or residual CO2 content required. It is good practice to operate the reboiler at as low a temperature as possible.
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Basics of Gas Field Processing Book
Prediction and Inhibition of Gas Hydrates Book
Basics of Corrosion in Oil and Gas Industry Book
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Chapter 6 216
Gas Sweetening 216
6.1 Introduction 216
6.1.1 Definition 216
6.1.2 Safety Precautions 216
6.1.3 Purification Levels 217
6.1.4 Acid Gas Disposal 217
6.1.5 Removal of contaminants and pretreatment 218
6.2 Gas Treating – Process Options 219
6.3 Chemical Solvent Processes 221
6.3.1 Amine (Aqueous Alkanolamine) Processes 223
6.3.2 Alkaline Salt Process (Hot Carbonate) 234
6.3.3 Specialty Batch Chemical Solvents 237
6.4 Physical Solvent Processes 238
6.4.1 Selexol® 239
6.4.2 Fluor Solvent 242
6.4.3 Rectisol Process® 243
6.4.5 Purisol® 243
6.4.5 Catasol® 243
6.4.6 Morphysorb® 243
6.5 Hybrid Solvent Processes 244
6.5.1 Sulfinol® Process 244
6.5.2 Selefining Process 244
6.6 General Considerations for Solvent Process 245
6.6.1 Solution Filtration 245
6.6.2 Flash Tank 245
6.6.3 Corrosion 246
6.6.4 Foaming 247
6.6.5 Materials 248
6.7 Solid Bed Processes 248
6.7.1 General Process Description 248
6.7.2 Iron Sponge Process 248
6.7.3 SulfaTreat® 250
6.7.4 Zinc Oxide Process 251
6.7.5 Chemsweet® 251
6.7.6 PuraSpec® 251
6.7.8 Molecular Sieve Process 251
6.7.9 Oxorbon 253
6.8 Direct Conversion Processes (Liquid Redox) 254
6.8.1 Stretford Process 255
6.8.2 Lo-Cat process 255
6.8.3 Sulferox process 257
6.8.4 IFP Process 258
6.9 Distillation Process 258
6.10 Sulfur Recovery (The Claus Process) 259
6.10.1 Claus Process Considerations 261
6.10.2 Process Variations 262
6.10.3 Combustion Operation 263
6.10.4 Claus Unit Tail Gas Handling 263
6.11 Gas Permeation Process (Membranes) 266
6.11.1 Membrane Fundamentals 266
6.11.2 Membrane Selection Parameters 266
6.11.3 Membrane Structure Types 267
6.11.4 Carbon Dioxide Removal from Natural Gas 268
6.11.5 Membrane Elements 268
6.11.6 Membrane Design Considerations 271
6.11.7 Operating Considerations 274
6.11.8 Feed Gas Pretreatment 277
6.11.9 Membrane Advantages & Disadvantages 280
6.11.10 Hybrid Configurations 282
6.12 Biological Processes 285
6.13 Process Selection 285
6.13.1 Inlet Gas Stream Analysis 285
6.13.2 General Considerations 285
6.13.3 Removal of H2S 286
6.13.4 Removal of H2S and CO2 286
6.13.5 Process Selection charts 287
6.14 Safety & Environmental Considerations 289
6.15 Design Procedure 290
6.15.1 Iron Sponge 290
Example 6-1 291
6.15.2 The Amine System 292
Method 1 292
Example 6-2 294
Method 2 294
Example 6-3 296
-----------------
Chapter 6
Gas Sweetening
6.1 Introduction
Gas treating involves reduction of the “acid gases” carbon dioxide (CO2) and hydrogen sulfide (H2S), along with other sulfur species, to sufficiently low levels to meet contractual specifications or permit additional processing in the plant without corrosion and plugging problems. This chapter focuses on acid gases because they are the most prevalent. When applicable, discussion is given to other sulfur species.
There are many methods that may be employed to remove acidic components (primarily H2S and CO2) and other impurities from hydrocarbon streams. The available methods may be broadly categorized as those depending on chemical reaction, absorption, adsorption or permeation. Processes employing each of these techniques will be described.
6.1.1 Definition
Acid Gases
H2S combined with water forms sulfuric acid. CO2 combined with water forms carbonic acid.
Both are undesirable because they cause corrosion and reduce heating value and sales value. H2S is poisonous and may be lethal.
Sour Gas:
Sour gas is defined as natural gas with H2S and other sulfur compounds.
Sweet Gas:
Sweet gas is defined as natural gas without H2S and other sulfur compounds.
Acid gases Partial Pressure:
Partial pressure is used as an indicator if treatment is required. Partial pressure is defined as
PP = (total pressure of system) x ( mol% of gas)
where CO2 is present with water, a partial pressure >30 psia, would indicate CO2 corrosion might be expected. Below 15 psia, would indicate CO2 corrosion would not normally be a problem although inhibition may be required.
Factors that influence CO2 corrosion are those directly related to solubility, that is, temperature, pressure, and composition of the water. Increased pressure increases solubility and increased temperature decreases solubility.
H2S may cause sulfide stress cracking due to hydrogen embrittlement in certain metals. H2S partial pressure >0.05 psia, necessitates treating.
6.1.2 Safety Precautions
Hydrogen sulfide is a highly toxic gas. At concentrations as low as 10 ppmv irritation of the eyes, nose, and throat is possible. The human nose can detect hydrogen sulfide in concentrations as low as 0.02 ppmv. However, the human sense of smell cannot be relied on to detect hazardous concentrations of hydrogen sulfide. Higher concentrations and extended exposure to hydrogen sulfide will desensitize the sense of smell. The concentrations required for different reactions by the human body are:
1. Threshold limit value (TLV) for prolonged exposure: 10 ppmv
2. Slight symptoms after several hours exposure: 10-100 ppmv
3. Maximum concentration that can be inhaled for one hour without serious effects such as significant eye and respiratory irritation: 200-300 ppmv
4. Dangerous after exposure of 30 minutes to one hour: 500-700 ppmv
5. Fatal in less than 30 minutes: 700-900 ppmv and above.
6. Death in minutes: greater than 1000 ppmv
Hydrogen sulfide is highly flammable and will combust in air at concentrations from 4.3 to 46.0 volume percent. Hydrogen sulfide vapors are heavier than air and may migrate considerable distances to a source of ignition.
When H2S concentrations are well above the ppmv level, other sulfur species can be present. These compounds include carbon disulfide (CS2), mercaptans (RSH), and sulfides (RSR), in addition to elemental sulfur. If CO2 is present as well, the gas may contain trace amounts of carbonyl sulfide (COS). The major source of COS typically is formation during regeneration of molecular-sieve beds used in dehydration (see Chapter 5).
Gaseous carbon dioxide is a naturally occurring gas that is 50% heavier than air and is colorless and odorless. It is also a principal by-product of combustion. CO2 is inactive and therefore non-flammable. CO2 will displace oxygen and can create an oxygen-deficient atmosphere resulting in suffocation. The principal hazard of CO2 is exposure to elevated concentrations.
The atmospheric concentration immediately hazardous to life is 10%. Because CO2 is heavier than air, its hazard potential is increased, especially when entering tanks and vessels.
"A common but erroneous belief is that CO2 simply acts as an asphyxiant by lowering the oxygen level below the 16 percent minimum necessary to sustain life (at sea level). Although this is frequently the case in most serious accidents, CO2 begins to have a noticeable effect on normal body functions at about two to three percent. The concentration of carbon dioxide in the blood affects the rate of breathing, a measurable increase resulting from a level of one percent in the inspired air."
6.1.3 Purification Levels
Although many natural gases are free of objectionable amounts of H2S and CO2, substantial quantities of these impurities are found in both gas reserves and production. In a survey of U.S. gas resources, Meyer (2000) defined subquality gas as that containing CO2 ≥ 2%, N2 ≥ 4%, or H2S ≥ 4 ppmv.
The inlet conditions at a gas processing plant are generally temperatures near ambient and pressures in the range of 300 to 1,000 psi (20 to 70 bar), so the partial pressures of the entering acid gases can be quite high. If the gas is to be purified to a level suitable for transportation in a pipeline and used as a residential or industrial fuel, then the H2S concentration must be reduced to 0.25 gr/100 scf (4 ppm, or 6 mg/m3), and the CO2 concentration must be reduced to a maximum of 3 to 4 mol%. However, if the gas is to be processed for NGL recovery or nitrogen rejection in a cryogenic turbo expander process, CO2 may have to be removed to prevent formation of solids. If the gas is being fed to an LNG liquefaction facility, then the maximum CO2 level is about 50 ppmv because of potential solids formation.
6.1.4 Acid Gas Disposal
What becomes of the CO2 and H2S after their separation from the natural gas?
The answer depends to a large extent on the quantity of the acid gases. For CO2, if the quantities are large, it is sometimes used as an injection fluid in EOR (enhanced oil recovery) projects. Several gas plants exist to support CO2 flooding projects. If this option is unavailable, then the gas can be vented, provided it satisfies environmental regulations for impurities.
In the case of H2S, four disposal options are available:
1. Incineration and venting, if environmental regulations regarding sulfur dioxide emissions can be satisfied
2. Reaction with H2S scavengers, such as iron sponge
3. Conversion to elemental sulfur by use of the Claus or similar process
4. Disposal by injection into a suitable underground formation.
The first two options are applicable to trace levels of H2S in the gas, and the last two are required if concentrations are too high to make the first two options feasible.
6.1.5 Removal of contaminants and pretreatment
All gas sweetening units should have well-designed pretreatment facilities. Carryover of brine or liquid hydrocarbon (as slugs or aerosol) from upstream production operations can cause problems for gas treating and downstream processing equipment. Also, field facilities are not typically designed to remove troublesome contaminants like gas-phase heavy hydrocarbons.
These contaminants can likewise cause operational difficulties in the sweetening process.
If gross liquid carryover from an upstream facility is possible, a slug catcher is recommended. It should be sized not only for steady inlet fluid volumes, but for surge capacity to handle slugs of liquid hydrocarbons, water, and/ or well treatment chemicals.
If aerosols are a concern, an inlet filter separator is suggested. Selected filter elements can remove entrained droplets down to 0.3 microns in diameter.
For liquid hydrocarbon treatment, a filter coalescer may be used to remove suspended water or glycol prior to further processing.
Heavy hydrocarbons (C6+) can be absorbed by solvents, which could lead to foaming in the sweetening unit. It is possible to reduce the heavy hydrocarbon content of the incoming gas through cooling (via Joule-Thomson expansion, propane refrigeration, or turbo-expansion), and subsequent condensation of the heavy components. The condensed liquids are removed, and the gas is warmed above the saturation temperature before going to the sweetening unit.
An alternative means to remove gaseous heavy hydrocarbons is through adsorption. Either alumina or silica gel beds may be used in parallel such that one bed is regenerated while the other is in service. The beds are regenerated by heating and desorbing the hydrocarbons. The heavy hydrocarbons are recovered from the regeneration gas via condensation.
Oxygen entry into a hydrocarbon system is often troublesome. If liquid water is present, severe corrosion may occur. If H2S or sulfur is present, corrosion by a different mechanism or sulfur deposition and plugging may occur. Oxygen contamination may be addressed by several different approaches but the first step is to find and correct the source of oxygen entry into the system. This is often the simplest and most cost effective approach. Most oxygen leaks may be traced to compressor suctions or pipe fittings.
To eliminate oxygen contamination a number of possibilities exist:
• React the oxygen with chemicals.
Chemicals such as amines, organics or inorganic compounds may be added to remove free oxygen. Oxygen scavengers are available from many chemical suppliers.
• Thermal oxidation reactions.
Integrated processes such as the DeOxy by “Optimized Process Designs” can perform a very limited burn to consume the free oxygen.
• Remove other reactants that cause problems with the presence of oxygen.
By removing offending components such as water or H2S that react with oxygen, the presence of low amounts of oxygen may be tolerated.
• Treat the symptom.
Corrosion inhibitors, filtration and / or alternate schemes may be utilized to stop or offset any adverse effects of oxygen contamination.
6.2 Gas Treating – Process Options
Sulfur exists in natural gas as hydrogen sulfide (H2S), and the gas is usually considered sour if the hydrogen sulfide content exceeds 4 ppm of H2S. The process for removing hydrogen sulfide and carbon dioxide from a natural gas stream is referred to as “sweetening” the gas.
Numerous processes have been developed for acid gas removal and gas sweetening based on a variety of chemical and physical principles.
Some of the more important items that must be considered before a process is selected are:
• The type and concentration of impurities and hydrocarbon composition of the sour gas. For example, COS, CS2, and mercaptans can affect the design of both gas and liquid treating facilities. Physical solvents tend to dissolve heavier hydrocarbons, and the presence of these heavier compounds in significant quantities tends to favor the selection of a chemical solvent.
• The temperature and pressure at which the sour gas is available. High partial pressures (50 psi [3.4 bar] or higher) of the acid gases in the feed favor physical solvents, whereas low partial pressures favor the amines.
• The specifications of the outlet gas (low outlet specifications favor the amines).
• The volume of gas to be processed.
• The specifications for the residue gas, the acid gas, and liquid products.
• The selectivity required for the acid gas removal.
• The capital, operating, and royalty costs for the process.
• The environmental constraints, including air pollution regulations and disposal of byproducts considered hazardous chemicals.
If gas sweetening is required offshore, both size and weight are additional factors that must be considered. Whereas CO2 removal is performed offshore, H2S removal is rarely done unless absolutely necessary because of the problems of handling the rich acid gas stream or elemental sulfur.
If the gas processing facility is to be used in conjunction with liquids recovery, the requirements for H2S, CO2, and mercaptan removal may be affected. In liquid recovery plants, varying amounts of H2S, CO2, and other sulfur compounds will end up in the liquid product. Failure to remove these components prior to liquids recovery may require liquid product treating in order to meet product specifications. In many instances, liquid treating may be required anyway.
When sulfur recovery is required, the composition of the acid gas stream feeding the sulfur plant must be considered. With CO2 concentrations greater than 80% in the acid gas, the possibility of selective treating should be considered to raise the H2S concentration to the sulfur recovery unit (SRU). This may involve a multi-stage gas treating system in which the gas exiting the first stage is enriched by passing it through another absorption solvent loop.
High concentrations of hydrocarbons can cause design and operating problems for the SRU. The effect of these components must be weighed when selecting the gas treating process to be used.
Decisions in selecting a gas treating process can be simplified by gas composition and operating conditions.
High partial pressures (50 psi) of acid gases enhance the possibility of using a physical solvent. The presence of significant quantities of heavy hydrocarbons in the feed discourages using physical solvents. Low partial pressures of acid gases and low outlet specifications generally require the use of amines for adequate treating. Process selection is not easy and a number of variables must be weighed prior to making a process selection.
Figure 6-1 summarizes the more important processes and groups them into the generally accepted categories.
Fig. 6-1. Acid Gas Removal Processes
Table 6.1 lists the processes used to separate the acid gas from other natural gas components.
Tables 6-2 and 6-3, include comparison of different gas sweetening processes.
Table 6.1 Acid Gas Removal Processes
Note a: MEA reacts nonreversibly with COS (carbonyl sulfide), and, therefore, should not be used to treat gases with a large concentration of COS.
Table 6.2 Gases Removed by Various Processes
6.3 Chemical Solvent Processes
Chemical reaction processes remove the H2S and/or CO2 from the gas stream by chemical reaction with a material in the solvent solution. The reactions may be reversible or irreversible.
In reversible reactions the reactive material removes CO2 and/or H2S in the contactor at high partial pressure and/or low temperature. The reaction is reversed by high temperature and/or low pressure in the stripper. In irreversible processes the chemical reaction is not reversed and removal of the H2S and/or CO2 requires continuous makeup of the reacting material.
Fig. 6-6 shows the process flow for a typical reversible chemical reaction process.
In solvent absorption, the two major cost factors are the solvent circulation rate, which affects both equipment size and operating costs, and the energy require and disadvantages of chemical and physical solvents. Table 6.4 summarizes some of the advantages and disadvantages of chemical and physical solvents.
Table 6-3. CO2 and H2S Removal Processes for Gas Streams.
The most common chemical solvents are:
• Amines (Aqueous Alkanolamine)
• Carbonates
Table 6.4. Comparison of Chemical and Physical Solvents
6.3.1 Amine (Aqueous Alkanolamine) Processes
All commonly used amines are alkanolamines, which are amines with OH groups attached to the hydrocarbon groups to reduce their volatility. Figure 6.2 shows the formulas for the common amines used in gas processing.
Amines are compounds formed from ammonia (NH3) by replacing one or more of the hydrogen atoms with another hydrocarbon group. Replacement of a single hydrogen produces a primary amine, replacement of two hydrogen atoms produces a secondary amine, and replacement of all three of the hydrogen atoms produces a tertiary amine. Primary amines form stronger bases than secondary amines, which form stronger bases than tertiary amines. Amines with stronger base properties are more reactive toward CO2 and H2S gases and form stronger chemical bonds. Sterically hindered amines are compounds in which the reactive center (the nitrogen) is partially shielded by neighboring groups so that larger molecules cannot easily approach and react with the nitrogen. The amines are used in water solutions in concentrations ranging from approximately 10 to 65 wt% amines.(table. 6-5 lists Amines physical properties ).
Table. 6-6. lists approximate guidelines for a number of alkanolamine processes.
The amine processes are particularly applicable where acid gas partial pressures are low and/or low levels of acid gas are desired in the residue gas.
Because the water content of the solution minimizes heavy hydrocarbon absorption, these processes are well suited for gases rich in heavier hydrocarbons. Some amines can be used to selectively remove H2S in the presence of CO2.
Amines remove H2S and CO2 in a two step process:
1. The gas dissolves in the liquid (physical absorption).
2. The dissolved gas, which is a weak acid, reacts with the weakly basic amines.
Absorption from the gas phase is governed by the partial pressure of the H2S and CO2 in the gas, whereas the reactions in the liquid phase are controlled by the reactivity of the dissolved species. The principal reactions are summarized in the next section.
Fig. 6.2 Molecular structures of commonly used amines.
Table 6-5 physical properties of gas sweetening chemicals.
Table 6-6. Approximate Guidelines for Amine Processes
Table. 6-5. Physical properties of gas treating chemicals.
Table 6-6. Approximate Guidelines for Amine Processes
Notes:
1. These data alone should not be used for specific design purposes. Many design factors must be considered for actual plant design.
2. Dependent upon acid gas partial pressures and solution concentrations.
3. Dependent upon acid gas partial pressures and corrosiveness of solution. Might be only 60% or less of value shown for corrosive systems.
4. Varies with stripper overhead reflux ratio. Low residual acid gas contents require more stripper trays and/or higher reflux ratios yielding larger reboiler duties.
5. Varies with stripper overhead reflux ratios, rich solution feed temperature to stripper and reboiler temperature.
6. Maximum point heat flux can reach 20,000–25,000 Btu/hr-ft2 at highest flame temperature at the inlet of a direct fired fire tube. The most satisfactory design of firetube heating elements employs a zone by zone calculation based on thermal efficiency desired and limiting the maximum tube wall temperature as required by the solution to prevent thermal degradation. The average heat flux, Q/A, is a result of these calculations.
7. Reclaimers are not used in DEA and MDEA systems.
8. Reboiler temperatures are dependent on solution conc. flare/vent line back pressure and/or residual CO2 content required. It is good practice to operate the reboiler at as low a temperature as possible.
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Gas Sweetening and Sulfur Recovery - Chapter 6 - Part 1 B.
6.3.1.1 Process Description
A typical amine system is shown in Figure 6.3. Sour gas enters the system through an inlet separator (scrubber) to remove any entrained water or hydrocarbon liquids. Gas enters the bottom of the amine absorber (contactor) and flows upward, countercurrent to the lean amine solution which flows down from the top. The lean amine that returns to the contactor is maintained at a temperature above the vapor that exits the contactor to prevent any condensation of heavier liquid hydrocarbons. Intimate contact between the gas and amine solution is achieved by use of either trays or packing in the contactor.
The absorber tower consists of trays (diameters >20 in. (500 mm)), conventional packing (diameters <20 in.) or structured packing (diameters >20 in.).
Sweetened gas leaves the top of the tower. An optional outlet (separator) scrubber may be included to recover entrained amine from the sweet gas. Since the natural gas leaving the top of the tower is saturated with water, the gas will require dehydration before entering a pipeline.
Rich amine, solution containing CO2 and H2S, leaves the bottom of the absorber and flows to the flash tank (drum) where most of the dissolved hydrocarbon gases or entrained hydrocarbon condensates are removed. A small amount of the acid gases flash to the vapor phase. From the flash drum, the rich amine proceeds to the rich amine/lean amine heat exchanger where it recovers some of the sensible heat from the lean amine stream, which decreases the heat duty on the amine reboiler and the solvent cooler. Preheated rich amine then enters the amine stripping tower where heat from the reboiler breaks the bonds between the amine and acid gases. Acid gases are removed overhead and lean amine is removed from the bottom of the stripper.
Hot lean amine flows to the rich amine/lean amine heat exchanger and then to additional coolers, typically aerial coolers, to lower its temperature about 100F (5.50C) above the inlet gas temperature. This reduces the amount of hydrocarbons condensed in the amine solution when the amine contacts the sour gas.
A side stream of amine, of about 3%, is taken off after the rich/lean amine heat exchanger and is flowed through a charcoal filter to clean the solution of contaminants (not included in figure). Cooled lean amine is then pumped up to the absorber pressure and enters the top of the absorber. Amine solution flows down the absorber where it absorbs the acid gases. Rich amine is then removed at the bottom of the tower and the cycle is repeated.
6.3.1.2 General Remarks
The contactor operates above ambient temperature because of the combined exothermic heat of absorption and reaction. The maximum temperature is in the lower portion of the tower because the majority of the absorption and reaction occurs near the bottom of the unit. The temperature “bulge” in the tower can be up to about 180°F (80°C). The treated gas leaves the top of the tower water saturated and at a temperature controlled by the temperature of the lean amine that enters, usually around 100°F (38°C).
The sweet gas leaving the contactor is saturated with water so dehydration, is normally required prior to sale. If MEA is the sweetening agent, or the contactor is operating at unusually high temperature, a water wash may be used to attempt to recover some of the vaporized and/or entrained amine from the gas leaving the contactor. If a water wash is used it generally will consist of three or four trays at the top of the contactor, with makeup water to the unit being used as the wash liquid.
Acid gas stripped from the amine passes out of the top of the stripper. It goes through a condenser and separator to cool the stream and recover water. The recovered water is usually returned to the stripper as reflux. The acid gas from the reflux separator is either vented, incinerated, sent to sulfur recovery facilities, or compressed for sale or reinjected into a suitable reservoir for enhanced oil recovery projects or for sequestration.
Fig. 6-3. Gas Sweetening by Amine absorption, operating conditions are representative, not definitive.
6.3.1.3 Operating Issues
Corrosion—Some of the major factors that affect corrosion are:
• Amine concentration (higher concentrations favor corrosion)
• Rich amine acid gas loading (higher gas loadings in the amine favor corrosion)
• Oxygen concentration
• Heat stable salts (higher concentrations promote corrosion and foaming)
In addition to destroying vessels and piping, the corrosion products can cause foaming.
Solution Foaming—Foaming of the liquid amine solution is a major problem because it results in poor vapor−liquid contact, poor solution distribution, and solution holdup with resulting carryover and off spec gas. Among the causes of foaming are suspended solids, liquid hydrocarbons, surface active agents, such as those contained in inhibitors and compressor oils, and amine degradation products, including heat stable salts. One obvious cure is to remove the offending materials; the other is to add antifoaming agents.
Heat Stable Salts (HSS) —As mentioned above, these amine degradation products can cause both corrosion and foaming. They are normally dealt with through the use of amine reclaimers.
6.3.1.4 Reclaimer
A reclaimer is usually required for MEA and DGA® amine-based systems. The reclaimer helps remove degradation products from the solution and also aids in the removal of heat stable salts, suspended solids, acids and iron compounds. The reclaimers in MEA and DGA® systems differ. For MEA, a basic solution helps reverse the reactions. Soda ash and/or caustic soda is added to the MEA reclaimer to provide a pH of approximately 8-9; no addition is required for the DGA® reclaimer system. Reclaimers generally operate on a side stream of 1-3% of the total amine circulation rate. Reclaimer sizing depends on the total inventory of the plant and the rate of degradation expected.
Reclaimer operation is a semi-continuous batch operation. The reclaimer is filled with hot amine solution and, if necessary, soda ash is added. As the temperature in the reclaimer increases, the liquid will begin to distill. Overhead vapors can be condensed and pumped back into the amine system, but generally the reclaimer is operated at slightly above stripper column pressure and the vapors are returned to the stripper.
The initial vapor composition is essentially water. Continued distillation will cause the solution to become more and more concentrated with amine. This raises the boiling point of the solution and amine will begin to distill overhead. Fresh feed is continually added until the boiling point of the material in the reclaimer reboiler reaches 280° to 300°F. At this point, distillation is continued for a short time adding only water to help recover residual amine in the reclaimer reboiler. The reclaimer is then cleaned, recharged, and the cycle is repeated.
Reclaimer "sludge" removed during cleaning must be handled with care. Disposal of the "sludge" must be in accordance with the governing regulations. If needed a reclaiming company may be contracted to remove degradation products or heat stable salts from the amine.
DEA does not form a significant amount of nonregenerable degradation products, and it requires more difficult reclaiming through vacuum distillation or ion exchange.
6.3.1.5 Basic Amine Chemistry
Amines are bases, and the important reaction in gas processing is the ability of the amine to form salts with the weak acids formed by H2S and CO2 in an aqueous solution. When a gas stream that contains the H2S, CO2, or both, is contacted by a primary or secondary amine solution, the acid gases react to form a soluble acid−base complex, a salt, in the treating solution. The reaction between the amine and both H2S and CO2 is highly exothermic. Regardless of the structure of the amine, the overall reactions between H2S and amines are simple since H2S reacts directly and rapidly with all amines to form the bisulfide and the sulfide by Eqs. 6-1 & 6-2:
For hydrogen sulfide removal
RNH2 + H2S ↔ RNH3+ + HS– Fast Eq. 6-1
RNH2 + HS– ↔ RNH3+ + S–2 Fast Eq. 6-2
For carbon dioxide removal
2RNH2 + CO2 ↔ RNH3+ + RNHCOO– Fast Eq. 6-3
RNH2 + CO2 + H2O ↔ RNH3+ + HCO3– Slow Eq. 6-4
RNH2 + HCO3– ↔ RNH3+ + CO3–2 Slow Eq. 6-5
Concerning the chemical reactions with CO2, primary amines (RNH2) such as MEA and DGA® agent, and secondary amines (RR’NH) such as DEA and DIPA, differ from tertiary amines (RR’R"N) such as TEA and MDEA.
For the reactions discussed above, high pressures and low temperatures drive the reactions to the right, whereas high temperatures and low pressures favor the reverse reaction, which thus provides a mechanism for regeneration of the amine solution.
Primary and Secondary Amines
With the primary and secondary amines, the predominant overall reaction (Eq. 6-3) rapidly leads to the formation of a stable carbamate which is slow to further hydrolyze to bicarbonate.
The other overall reactions leading to bicarbonate (Eq. 6-4) and to carbonate (Eq. 6-5) are slow because they have to proceed through the hydration of CO2.
Therefore, according to Eq. 6-3 there is a theoretical limit to the chemical loading capacity of the primary and secondary amine solutions to 0.5 mole CO2 per mole of amine, even at relatively high partial pressures of CO2 in the gas to be treated.
For primary and secondary amines, little difference exists between the H2S and CO2 reaction rates because of the availability of the rapid carbamate formation for CO2 absorption. Therefore, the primary and secondary amines achieve essentially complete removal of H2S and CO2.
Tertiary Amines
Unlike primary and secondary amines, the nitrogen (N) in tertiary amines (RR’R"N) has no free hydrogen ( H ) to rapidly form carbamate as per overall Eq. 6-3. Consequently, the removal of CO2 by tertiary amines can only follow the slow route to bicarbonate by Eq. 6-4 and carbonate by Eq. 6-5.
The slowness of the reaction leading to bicarbonate is the underlying reason why tertiary amines can be considered selective for H2S removal, by playing with absorption contact time, and this attribute can be used to full advantage when complete CO2 removal is not necessary.
However, the slow route to bicarbonates theoretically allows at equilibrium a chemical loading ratio of one mole of CO2 per mole of amine. Furthermore, at high partial pressure, the solubility of CO2 in tertiary amines is far greater than in the primary and secondary amines thus further enhancing the CO2 loading by physical solubility at high partial pressure.
Therefore, in case of gases to be treated for bulk CO2 removal, large amounts of CO2 can be liberated from the rich solvent by simple flash alleviating the thermal regeneration duty with consequent energy savings. In other words, because the tertiary amines have no labile hydrogen, they cannot form the carbamate. Tertiary amines must react with CO2 via the slow hydrolysis mechanism in Equation 6.4. With only the slow acid−base reaction available for CO2 absorption, MDEA (methyldiethanolamine) and several of the formulated MDEA products yield significant selectivity toward H2S relative to CO2, and, consequently, all of the H2S is removed while some of the CO2 “slips” through with the gas.
Activated Tertiary Amines
The use of activators mitigates the slowness of the reaction to bicarbonate for tertiary amines. Activators are generally primary or secondary amines; they are tailored to increase both the hydrolysis of the carbamate and the rate of hydration of dissolved CO2 thus making the activated-tertiary amines specially suitable for efficient and economical bulk CO2 removal when selectivity is not required (see section on MDEA).
6.3.1.6 Amines Used
Monoethanolamine
MEA is a primary amine, which has had widespread use as a gas sweetening agent. The process is well proven and can meet pipeline specifications. MEA is a stable compound and, in the absence of other chemicals, suffers no degradation or decomposition at temperatures up to its normal boiling point.
Gas sweetening with monoethanolamine (MEA) is used where there are low contactor pressures and/or stringent acid gas specifications. MEA removes both H2S and CO2 from gas streams. H2S concentrations well below 4.0 ppmv can be achieved. CO2 concentrations as low as 100 ppmv can be obtained at low to moderate pressures.
COS and CS2 are removed by MEA, but the reactions are irreversible unless a reclaimer is used. Even with a reclaimer, complete reversal of the reactions may not be achieved. The result is solution loss and build-up of degradation products in the system. Total acid gas pick up is traditionally limited to 0.3-0.35 moles of acid gas/mole of MEA and solution concentration is usually limited to 10-20 wt%. Inhibitors can be used to allow much higher solution strengths and acid gas loadings. Because MEA has the highest vapor pressure of the amines used for gas treating, solution losses through vaporization from the contactor and stripper can be high (MEA losses of 1-3 lbs/MMscf (16-48 kg/ MM m3) of inlet gas). This problem can be minimized by using a water wash.
MEA reactions are reversible by changing the system temperature. Reactions with CO2 and H2S are reversed in the stripping column by heating the rich MEA to about 2450F at 10 psig (118 0C at 69 kPa). Acid gases evolve into the vapor and are removed from the still overhead. Thus, the MEA is regenerated.
MEA systems foam rather easily resulting in excessive amine carryover from the absorber. Foaming can be caused by a number of foreign materials such as condensed hydrocarbons, degradation products, solids such as carbon or iron sulfide, excess corrosion inhibitor, and valve grease.
A microfiber filter separator should be installed at the gas inlet to the MEA contactor. It is an effective method of foam control and removes many of the contaminants before they enter the system. Hydrocarbon liquids are usually removed in the flash tank. Degradation products are removed in a reclaimer as described above.
A gas blanket system is installed on MEA storage tanks and surge vessels. This prevents oxidation of MEA. Sweet natural gas or nitrogen is normally used.
Disadvantages include the reaction of MEA with carbonyl sulfide (COS) and carbon disulfide (CS2) to form heat-stable salts, which cannot be regenerated at normal stripping column temperatures. At temperatures above 245 0F (118 0C) a side reaction with CO2 exists that produces oxazolidone-2, a heat-stable salt, which consumes MEA from the process. Normal regeneration temperature in the still will not regenerate heat-stable salts or oxazolidone-2. A reclaimer is often included to remove these contaminants.
Diethanolamine
This process employs an aqueous solution of diethanolamine (DEA) (a secondary amine). DEA will not treat to pipeline quality gas specifications at as low a pressure as will MEA.
This process is used for high pressure, high acid gas content streams having a relatively high ratio of H2S/CO2.
Although mole/mole loadings as high as 0.8-0.9 have been reported, most conventional DEA plants still operate at significantly lower loadings.
The process flow scheme for conventional DEA plants resembles the MEA process. The advantages and disadvantages of DEA as compared to MEA are:
• The mole/mole loadings typically used with DEA (0.35-0.82 mole/mole) are much higher than those normally used (0.3-0.4) for MEA.
• Molecular weight of DEA is 105 compared to 61 for MEA. Combination of molecular weights and reaction stoichiometry means that about 1.7 lbs (0.77 kg) of DEA must be circulated to react with the same amount of acid gas as 1.0 lbs (0.45 kg) of MEA. The solution strength of DEA ranges up to 35% by weight compared to 20% for MEA. Loadings for DEA systems range from 0.35 to 0.65 mol of acid gas per mole of DEA without excessive corrosion. The result of this is that the circulation rate of a DEA solution is slightly less than in a comparable MEA system.
• Because DEA does not form a significant amount of nonregenerable degradation products, a reclaimer is not required.
• DEA is a secondary amine and is chemically weaker than MEA, and less heat is required to strip the amine solution.
• DEA forms a regenerable compound with COS and CS2 and can be used for the partial removal of COS and CS2 without significant solution losses.
• As a secondary amine, DEA is less alkaline than MEA. DEA systems do suffer the same corrosion problems, but not as severely as those using MEA. Solution strengths are typically from 25% to 35% DEA by weight in water.
• Vapor pressure of DEA is about 1/30 of the vapor pressure of MEA. Thus, DEA amine losses are much lower than in an MEA system.
Diglycolamine
DGA® is a primary amine capable of removing not only H2S and CO2, but also COS and mercaptans from gas and liquid streams. DGA® has been used to treat natural gas to 4.0 ppmv at pressures as low as 125 psig. Compared with MEA, low vapor pressure allows Diglycolamine [ 2-(2-aminoethoxy) ethanol] (DGA) is reclaimed on-site to remove heat stable salts and reaction products with COS and CS2.
DGA® has a greater affinity for the absorption of aromatics, olefins, and heavy hydrocarbons than the MEA and DEA systems.
Therefore, adequate carbon filtration should be included in the design of a DGA® treating unit.
The process flow for the DGA® treating process is similar to that of the MEA treating process. The major differences are:
• Higher acid gas pick-up per gallon of amine can be obtained by using 50-60% solution strength rather than 15-20% for MEA (more moles of amine per volume of solution).
• The required treating circulation rate is lower. This is a direct function of higher amine concentration.
• Reduced reboiler steam consumption.
• DGA® has an advantage for plants operating in cold climates where freezing of the solution could occur. The freezing point for 50% DGA® solution is –30°F.
• Low vapor pressure decreases amine losses.
• Unlike MEA, degradation products from reactions with COS and CS2 can be regenerated.
Methyldiethanolamine
Methyldiethanolamine (MDEA) is a tertiary amine which can be used to selectively remove
H2S to pipeline specifications at moderate to high pressure. If increased concentration of CO2 in the residue gas does cause a problem with contract specifications or downstream processing, further treatment will be required.
If the gas is contacted at pressures ranging from 800 to 1000 psig, H2S levels can be reduced to concentrations required by pipelines. While at the same time, 40-60% of the CO2 present flows through the contactor, untreated.
The H2S/CO2 ratio in the acid gas can be 10-15 times as great as the H2S/CO2 ratio in the sour gas. Some of the benefits of selective removal of H2S include:
• Reduced solution flow rates resulting from a reduction in the amount of acid gas removed.
• Smaller amine regeneration unit. Significant capital savings are realized due to reduced pump and regeneration requirements.
• Higher H2S concentrations in the acid gas resulting in reduced problems in sulfur recovery.
• MDEA has a lower heat requirement due to its low heat of regeneration. In some applications, energy requirements for gas treating can be reduced as much as 75% by changing from DEA to MDEA.
• It is not reclaimable by conventional methods.
CO2 hydrolyzes much slower than H2S. This makes possible significant selectivity of tertiary amines for H2S. This fact is used for selective removal of H2S from gases containing both H2S and CO2.
A feature of MDEA is that it can be partially regenerated in a simple flash. As a consequence the removal of bulk H2S and CO2 may be achieved with a modest heat input for regeneration.
However, as MDEA solutions react only slowly with CO2, activators must be added to the MDEA solution to enhance CO2 absorption and the solvent is then called activated MDEA.
Solution strengths typically range from 40% to 50% MDEA by weight. Acid gas loading varies from 0.2 to 0.4 or more moles of acid gas per mole of MDEA depending on the supplier.
Diisopropanolamine
Diisopropanolamine (DIPA) is a secondary amine which exhibits, though not as great as tertiary amines, selectivity for H2S. It is similar to DEA systems but offers the following advantages:
• Carbonyl sulfide (COS) can be removed and the DIPA solution can be regenerated easily
• The system is generally noncorrosive and has a lower energy consumption.
• At low pressures, DIPA will preferentially remove H2S. As pressure increases, the selectivity of the process decreases and DIPA removes increasing amounts of CO2. Thus, this system can be used either to selectively remove H2S or to remove both CO2 and H2S.
Formulated Solvents and Mixed Amines
The selectivity of MDEA can be reduced by addition of MEA, DEA, or proprietary additives. Thus, it can be tailored to meet the desired amount of CO2 slippage and still have lower energy requirements than do primary and secondary amines.
Formulated Solvents is the name given to a new family of amine based solvents. Their popularity is primarily due to equipment size reduction and energy savings over most of the other amines. All the advantages of MDEA are valid for the Formulated Solvents, usually to a greater degree. Some formulations are capable of slipping larger portions of inlet CO2 (than MDEA) to the outlet gas and at the same time removing H2S to less than 4 ppmv. For example, under conditions of low absorber pressure and high CO2 /H2S ratios, such as Claus tail gas clean-up units, certain solvent formulations can slip upwards to 90 percent of the incoming CO2 to the incinerator.
While at the other extreme, certain formulations remove CO2 to a level suitable for cryogenic plant feed. Formulations are also available for CO2 removal in ammonia plants. Finally, there are solvent formulations which produce H2S to 4 ppmv pipeline specifications, while reducing high inlet CO2 concentrations to 2% for delivery to a pipeline. This case is sometimes referred to as bulk CO2 removal.
This need for a wide performance spectrum has led “Formulated Solvent” suppliers to develop a large stable of different MDEA-based solvent formulations. Most Formulated Solvents are enhancements to MDEA discussed above. Thus, they are referred to as MDEA-based solvents or formulations.
Benefits claimed by suppliers are:
For New Plants
• reduced corrosion
• reduced circulation rate
• lower energy requirements
• smaller equipment due to reduced circulation rates
For Existing Plants
• increase in capacity, i.e., gas through-put or higher inlet acid gas composition
• reduced corrosion
• lower energy requirements and reduced circulation rates
Inhibited Amine Systems
Inhibited amine processes use standard amines that have been combined with special inhibiting agents that minimize corrosion. They allow higher solution concentrations and higher acid gas loadings, thus reducing required circulation rates and energy requirements. They also utilize hot potassium carbonate to remove CO2 and H2S. As a general rule, this process should be considered when the partial pressure of the acid gas is 20 psia (138 kPa) or greater. The process is not recommended for low-pressure absorption, or high-pressure absorption of low-concentration acid gas.
Sterically Hindered Amines
In acid gas removal, steric hindrance involves alteration of the reactivity of a primary or secondary amine by a change in the alkanol structure of the amine.
A large hydrocarbon group attached to the nitrogen shields the nitrogen atom and hinders (inhibits) the carbamate reaction. The H2S reaction is not significantly affected by amine structure, because the proton is small and can reach the nitrogen. However, CO2 removal can be significantly affected if the amine structure hinders the fast carbamate formation reaction and allows only the much slower bicarbonate formation.
Table 6.7. Some Representative Operating Parameters for Amine Systems
6.3.1.7 Heats of Reaction
Amines are basic solutions. These solutions react with the hydrogen sulfide and carbon dioxide to form a salt. The process of absorbing the acid gases generates heat.
The magnitude of the exothermic heats of reaction, which includes the heat of solution, of the amines with the acid gases is important because the heat liberated in the reaction must be added back in the regeneration step. Thus, a low heat of reaction translates into smaller energy regeneration requirements. Table 6.8 summarizes the important data for the common amines. The values in Table 6.8 are approximate because heats of reaction vary with acid gas loading and solution concentration.
Table 6.8. Average Heats of Reaction of the Acid Gases in Amine Solutions
6.3.2 Alkaline Salt Process (Hot Carbonate)
The hot potassium carbonate process for removing CO2 and H2S was developed by the United States Bureau of Mines and is described by Benson and coworkers in two papers. Although the process was developed for the removal of CO2, it can also remove H2S if H2S is present with CO2 (not suitable for gas streams containing only H2S.). Special designs are required for removing H2S to pipeline specifications or to reduce CO2 to low levels.
The process is very similar in concept to the amine process, in that after physical absorption into the liquid, the CO2 and H2S react chemically with the solution. The chemistry is relatively complex, but the overall reactions are represented by
K2CO3 + CO2 + H2O ↔ 2KHCO3 Eq. 6-6
K2CO3 + H2S ↔ KHS + KHCO3 Eq. 6-7
In a typical application, the contactor will operate at approximately 300 psig (20 barg), with the lean carbonate solution entering near 225°F (110°C) and leaving at 240°F (115°C). The rich carbonate pressure is reduced to approximately 5 psig (0.3 barg) as it enters the stripper.
Approximately one third to two thirds of the absorbed CO2 is released by the pressure reduction, reducing the amount of steam required for stripping. The lean carbonate solution leaves the stripper at the same temperature as it enters the contactor, and eliminates the need for heat exchange between the rich and lean streams. The heat of solution for absorption of CO2 in potassium carbonate is small, approximately 32 Btu/cu ft of CO2, and consequently the temperature rise in the contactor is small and less energy is required for regeneration.
Potassium carbonate processes are somewhat effective in removing carbonyl sulfide and carbon disulfide. Potassium carbonate works best on gas streams with a CO2 partial pressure of 30-90 psi.
Pipeline-quality gas often requires secondary treating using an amine or similar system to reduce H2S level to 4 ppm.
Because this system is operated at high temperatures to increase the solubility of carbonates, the designer must be careful to avoid dead spots in the system where the solution could cool and precipitate solids. If solids do precipitate, the system may suffer from plugging, erosion, or foaming.
Hot potassium carbonate solutions are corrosive. All carbon steel must be stress relieved to limit corrosion. Varieties of corrosion inhibitors, such as fatty amines or potassium dichromate, are available to decrease corrosion rates. There are three basic process flow variations for the potassium carbonate process. The flow scheme required depends on the outlet specification of the natural gas. These are:
Single Stage Process
The single stage process is shown in Fig. 6-4. Potassium carbonate is pumped to the top of a packed or trayed contactor where it contacts the gas stream. The rich solution flows to the stripper where the acid gases are stripped with steam. The lean solution is then pumped back to the contactor to complete the cycle.
Split Flow Process
In this process scheme (Fig. 6-5) the lean solution stream is split. Hot solution is fed to the middle of the contactor for bulk removal. The remainder is cooled to improve equilibrium and is fed to the top of the contactor for trim acid gas removal.
Two Stage Process
In this process scheme (Fig. 6-6) the contactor is like that of the split flow process. In addition, the stripper is in two sections. A major portion of the solution is removed at the midpoint of the stripper and pumped to the lower section of the contactor.
Numerous improvements have been made to the potassium carbonate process resulting in significant reduction in capital and operating costs. At the same time, lower acid gas concentration in the treated gas can now be achieved. The most popular of the carbonate processes are:
Benfield Process
The Benfield Process is licensed by UOP. Several activators are used to enhance the performance of the potassium carbonate solution.
Hi-Pure Process
The Hi-Pure process is a combination conventional Benfield potassium carbonate process and alkanolamine process. The gas stream is first contacted with potassium carbonate followed by contacting with an amine. The process can achieve outlet CO2 concentrations as low as 30 ppmv and H2S concentrations of 1 ppmv.
Catacarb Process
The Catacarb Process is licensed by Eickmeyer and Associates. Activators, corrosion inhibitors, potassium salts, and water are contained in the solution. This process is mostly used in the ammonia industry.
Fig. 6-4. Alkaline Salt: Single-Stage Process
Fig. 6-5. Alkaline Salt: Split-Flow Process
Fig. 6-6. Alkaline Salt: Two-Stage Process
6.3.3 Specialty Batch Chemical Solvents
Several batch chemical processes have been developed and have specific areas of application.
Processes include
General Process Description
Gas is flowed into a vessel and contacted with the solvent. Acid components are converted to soluble salts, which are nonregenerable, limiting the life of the solution.
Once saturation levels are reached, the solution must be replaced.
For some of these processes, the spent solutions are not hazardous, but for others, the spent solutions have been labeled hazardous and, if used, must be disposed of as Class IV materials.
Units in these processes have a wide operating range, with acid gas concentrations ranging from as low as 10 ppm to as high as 20%. Operating pressures range from near atmospheric to >1000 psig (7000 kPa). Some units have been designed to handle from several thousand cubic feet per day to more than 15 MMscfd (several hundred cubic meters per day to more than 420,000 m3 per day).
a- Sulfa-Check
Sulfa-Check is a product of ExxonMobil Chemicals which selectively removes H2S and mercaptans from natural gas in the presence of CO2. The patented process converts the sour gas directly to sulfur. This is accomplished by sparging “bubbling” the gas in a buffered, water based oxidizing solution containing sodium nitrite (NaNO2). The sodium nitrite is reduced to ammonia (NH3) which remains in solution. The spent product is classified as non-hazardous.
Dozens of field applications include gas rates ranging from 15 Mscf/d to 3.0 MMscf/d and inlet H2S concentrations ranging from 10 to 3000 ppmv.
Reaction rate is independent of the concentration of the oxidizing agent.
There is no limit to the concentration of H2S treated. Process is most economical for acid gas streams containing from1 ppm to 1% H2S. pH must be held above 7.5 to control selectivity and optimize H2S removal. One gallon (4 l) of oxidizing solution can remove up to 2 lbs (1 kg) of H2S when the system is operated at ambient temperatures<100 0F (38 0C). If gas temperatures exceed 100 0F (38 0C), the solubility of sulfur in the oxidizing agent decreases.
Operating pressure of at least 20 psig is required for proper unit operation to maintain bubble flow through the column. Bubble flow is necessary to produce intimate mixing of the gas and liquid.
Oxidizing solution will eventually become saturated and require replacement. Disposal of this slurry poses no environmental problem, as the reaction produces an aqueous slurry of sulfur and sodium salt.
A number of variables, including some associated risks must be considered prior to determining if the Sulfa-Check process is applicable. For example, low levels of ammonia may appear in the treated gas. Also, the reduction of NO2 may result in the formation of nitric oxide (NO). If air is present in the raw gas, it will react with nitric oxide to form nitrogen dioxide (NO2). NO2 is a strong oxidizing agent that will react with elastomers and odorants and cause corrosion in a moist environment.
b- Caustic Wash
Caustic (NaOH) scrubbing systems can be used to treat natural gas streams to remove CO2, CS2, H2S, and mercaptans. The process employs countercurrent contacting of the gas stream with a caustic solution in a packed or trayed column. The column may contain one stage or several stages depending on the required degree of removal. The multi-stage systems generally have different caustic concentrations ranging from 4-6 weight percent in the first stage to 8-10 weight percent in the latter stages. Multiple stages increase the caustic efficiency while maintaining a sufficient driving force to achieve absorption.
The spent solution is either regenerated or discarded depending on what acid gas components are present in the gas stream. If only mercaptans are present, the caustic solution is regenerated with steam in a stripping still. If CO2 is present, a nonregenerable product (Na2CO3) is formed and the solution must be discarded. As a result, the presence of CO2 in caustic systems leads to high caustic consumption. This is a serious disadvantage of the caustic scrubbing process. The spent caustic solutions are considered hazardous wastes. Natural gas is usually water washed after a caustic wash to remove any caustic entrained in the gas prior to dehydration.
The chemical reactions involved are as follows:
H2S + 2 NaOH → Na2S + 2H2O Eq 6-8
RSH + NaOH → RSNa + H2O Eq 6-9
CO2 + 2 NaOH → Na2CO3 + H2O Eq 6-10
CS2 + 2 NaOH → 2 NaHS + CO2 Eq 6-11
Fig. 6-7. Regenerative Caustic
c- Sulfa-Scrub
Sulfa-Scrub from Baker-Petrolite is a product that uses triazine compound to selectively react with H2S. The raw gas can be sparged into a vessel containing the liquid scavenging agent. Alternatively the liquid may be injected to the gas stream as an H2S scavenger. The spent material is considered as non-hazardous and is claimed to be an excellent corrosion inhibitor.
6.4 Physical Solvent Processes
In the amine and alkali salt processes, the acid gases are removed in two steps:
physical absorption followed by chemical reaction. In processes such as Selexol® or Rectisol®, no chemical reaction occurs and acid gas removal depends entirely on physical absorption. Some of the inherent advantages and disadvantages of physical absorption processes are summarized below:
These processes flow scheme is as shown in Fig. 6-8. In general, a physical solvent process should be considered when:
1- The partial pressure of the acid gas in the feed is greater than 50 psi.
2- The heavy hydrocarbon concentration in the feed gas is low.
3- Bulk removal of the acid gas is desired.
4- Selective removal of H2S is desired.
These processes are economically attractive because little energy is required for regeneration. The solvents are regenerated by:
1- Multi-stage flashing to low pressures.
2- Regeneration at low temperatures with an inert stripping gas.
3- Heating and stripping of solution with steam/solvent vapors.
In general, physical solvents are capable of removing COS, CS2, and mercaptans.
In certain instances, physical absorption processes are capable of simultaneously dehydrating and treating the gas although additional equipment and higher energy requirements may be needed to dry the solvent. The processes operate at ambient or subambient temperature to enhance the solubility of the acid gases. The solvents are relatively noncorrosive so carbon steel can be used. Chemical losses are low due to low solvent vapor pressure or refrigerated conditions. Physical solvents will absorb heavy hydrocarbons from the gas stream resulting in high hydrocarbon content in the acid gas stream as well as possibly significant hydrocarbon losses. Some of the physical absorption processes are summarized below.
Fig. 6-8. Typical Gas Sweetening by Physical Absorption
6.4.1 Selexol®
This process developed by Allied Chemical Corp. uses a polyethylene glycol derivative as a solvent. The solvent is selective for RSH, CS2, H2S, and other sulfur compounds. The process can be used to selectively or simultaneously remove sulfur compounds, carbon dioxide, water, as well as paraffinic, olefinic, aromatic and chlorinated hydrocarbons from a gas or air stream. Because water and heavy hydrocarbons are highly soluble in Selexol®, the treated gas from a Selexol® unit normally meets both water and hydrocarbon dew point specifications.
Levels of CO2 can be reduced by about 85%. The process is economical when high acid gas partial pressures exist and there is an absence of heavy ends in the gas.
This process will not typically remove enough CO2 to meet pipeline gas requirements. DIPA can be added to the solution to remove CO2 down to pipeline specifications. The process also removes water to <7 lbs/MMscf (0.11 g/std m3). The addition of DIPA increases the relatively low stripper heat duty.
The vendor states that the solvent is very stable, no degradation products are formed or disposed of, and no solvent reclaiming is required. Depending on the applications, the operating pressure could be as low as ambient though higher pressure is preferred. Operating temperature varies from 0°F to ambient. Selexol® is a Dow Chemical solvent and a UOP technology.
Table 6-9 presents RK, the ratio of the K-value* for methane, KCH4 (arbitrarily assigned a value of 1), to the K-values of the other component, RK = KCH4/Kcomponent . (The K-value is the ratio of the mole fraction of the component in the vapor phase (y) to its mole fraction in the liquid phase (x), K = y/x. High K-values indicate the material is predominately in the vapor phase, whereas low K-values indicate a higher concentration in the liquid phase (x)).
Table 6-9 Typical Relative Ratio of K-values
An R value greater than1 indicates that the solubility of the component in Selexol® is greater than that of methane. The values should be regarded as only representative because pressure and temperature are not specified and, as previously noted, the composition of Selexol® is variable.
Because RK for CO2 and H2S are 15 and 134, respectively, these gases are preferentially absorbed (relative to CH4), and, consequently, physical absorption is an effective technique for acid gas removal. The process can reduce H2S to 4 ppmv, reduce CO2 to levels below 50 ppmv, and essentially remove all mercaptans, CS2, and COS. Two additional features of Table 6-9 are worth mentioning. Because the RK values for hydrocarbons heavier than CH4 are fairly high (6.4 for C2H6, 15.3 for C3H8, and 35 for n-C4H10), Selexol® will remove substantial quantities of these hydrocarbons, a feature that can be either positive or negative, depending on the composition of the gas being processed and the desired products. Finally, the RK value of H2O is extremely high and consequently, Selexol® provides some dehydration.
Volumes of some gases in scf absorbed/gal Selexol® are plotted in Figure 6-9. The figure assumes a Henry’s law relationship, (For an ideal system, Henry’s law assumes a linear relation between the solubility of gas component (i) and its partial pressure, yiP = kixi where ki is the Henry’s constant.) which provides approximate solubility at higher pressures. The lines also ignore probable Interaction between solutes.
Applications of Selexol® are varied and, consequently, no common process flow diagrams are available. One plant is shown in Figure 6-10. The plant pretreats the gas to reduce CO2, ethane, and heavier hydrocarbon levels before final purification in molecular sieve units and subsequent liquefaction. The plant is designed to process 26 MMscfd (0.74 MMSm3/d) entering the Selexol® unit at 603 psia (41.6 bar) and 32°F (0°C). The lean solvent, cooled to 25°F (−3.9°C) with propane refrigerant, enters the absorber where it absorbs CO2 and some of the ethane and heavier hydrocarbons. The rich solvent from the absorber is regenerated by reduction of the pressure in three flash drums, from 603 to 106 psia (41.6 to 7.3 bar) in the high-pressure drum, from 106 to 16 psia (7.3 to 1.1 bar) in the medium-pressure drum, and from 16 to 3 psia (1.1 to 0.21 bar) in the vacuum drum. Lean Selexol® from the vacuum drum is recompressed and sent to the propane chiller. The treated gas that leaves the absorber passes through a knockout drum and filter separator to remove entrained Selexol® and condensed hydrocarbons. Table 6-11 shows that the treated gas meets the specifications of a maximum of 0.50% CO2 and a maximum of 6.5% ethane and heavier hydrocarbons. In addition, the water content of the gas is reduced from 75 ppmv to 12 ppmv, H2S is reduced from 2 ppmv to essentially nothing, and methyl mercaptan is reduced from 5 ppmv to 1 ppmv. Unlike the amine systems, no irreversible products are generated in the process, which thus eliminates the need for reclaiming.
Table 6-10. Representative Property Data for Selexol®
Fig. 6-9. Solubility of various gases in Selexol® solvent at 70°F (21°C) as a function of partial pressure.
Table 6-11. Composition of Inlet and Outlet Gas in a Selexol® Unit
Fig. 6-10. Process schematic for a Selexol® gas treating facility.
6.4.2 Fluor Solvent
The Fluor Solvent process uses propylene carbonate as a physical solvent to remove CO2 and H2S. Propylene carbonate also removes C3+ hydrocarbons, COS, SO2, CS2, and H2O from the natural gas stream. Thus, in one step the natural gas can be sweetened and dehydrated to pipeline quality.
This process is used for bulk removal of CO2 and is not used to treat to <3% CO2.
This system requires special design features such as larger absorbers and higher circulation rates to obtain pipeline quality and usually is not economically applicable for outlet requirements.
Propylene carbonate has the following characteristics, which make it suitable as a solvent for acid gas treating:
• High degree of solubility for CO2 and other gases
• Low heat of solution for CO2
• Low vapor pressure at operating temperature
• Low solubility for light hydrocarbons (C1, C2)
• Chemically nonreactive toward all natural gas components
• Low viscosity
• Noncorrosive toward common metals
The above characteristics combine to yield a system that has low heat and pumping requirements, is relatively noncorrosive, and suffers only minimal solvent losses (<1 lbs/MMscf).
Solvent temperatures below ambient are often used to increase solvent gas capacity and, therefore, decrease circulation rates. Expansion of the rich solvent and flash gases through power turbines can provide the required refrigeration. Alternately, auxiliary refrigeration may be included to further decrease circulation rates.
6.4.3 Rectisol Process®
This process uses pure refrigerated methanol as a solvent and has been developed and licensed by Lurgi Oel Gas Chemie and Linde. The process is often applied for syngas purification and operates at temperatures as low as minus 30 to minus 100°F. Applications include selective or bulk removal of sour components. However the extremely low total sulfur impurities achievable with this process (down to less than 0.1 ppmv), are normally not required in natural gas service. Here the process is best suited where there are very high amounts of CO2 to be removed.
The process has been applied for the purification of natural gas for LNG production. However as pure cold methanol exhibits a certain co-absorption of higher hydrocarbons, implementation of the process in natural gas service has been limited to applications where only low concentrations of ethane and heavier components are present.
6.4.4 Ifpexol Process®
This process developed by IFP licensed by Prosernat uses refrigerated methanol solutions containing water in order to reduce hydrocarbon co-absorption. It is generally associated downstream of the Ifpex-1 dehydration process (Dehydration chapter) which simultaneously extracts heavier hydrocarbons. The Ifpex-2 process attains equivalent acid gas and mercaptans removal performance of the Rectisol process but at a slightly higher solvent rate due to the lower solvent purity.
6.4.5 Purisol®
This process has been developed and licensed by Lurgi Oel Gas Chemie. The solvent used is N-methyl-2-pyrrolidone (NMP or N-Pyrol), a high boiling point liquid. Purisol exhibits a selectivity for H2S, like Selexol, and features equivalent process possibilities.
As Purisol is also well suited for the selective removal of mercaptans; it can be used for the recovery of mercaptans from regeneration off gases from adsorptive mercaptan removal units.
6.4.5 Catasol®
This process is licensed by Eickmeyer and Associates. Catasol solvents are reported to have selectivity for H2S/CO2, CO2/propane, and COS/CO2.
6.4.6 Morphysorb®
A physical solvent process developed by Krupp Uhde called Morphysorb, uses N-formyl morpholine (NFM) as a physical solvent. The process claims lower investment and operating costs when processing gases heavily loaded with acid components.
6.5 Hybrid Solvent Processes
Table 6-4 showed the strengths and weaknesses of amine and physical solvent systems. To take advantage of the strengths of each type, a number of hybrid processes commercially used, and under development, combine physical solvents with amines. Depending upon the solvent−amine combination, nearly complete removal of H2S, CO2, and COS is possible. Other hybrid systems provide high H2S and COS removal while slipping CO2. Sulfinol® currently is one of the more commonly used processes. The process uses a combination of a physical solvent (sulfolane) with DIPA or MDEA. The selected amine depends upon the acid gases in the feed and whether CO2 removal is required. Like the physical solvent processes, the hybrid systems may absorb more hydrocarbons, including BTEX, but that property can be adjusted by varying water content.
6.5.1 Sulfinol® Process
The Sulfinol® Process, licensed by Shell Global Solutions, is used to remove H2S, CO2, COS, CS2, mercaptans and polysulfides from natural and synthetic gases. Sulfinol is a mixture of Sulfolane (a physical solvent), water and either DIPA or MDEA (both chemical solvents).
Solution concentrations range between 25% and 40% sulfolane, 40-55% DIPA, and 20-30% water and depend on the conditions and composition of the gas being treated. It is this dual capacity as both a physical and a chemical solvent that gives Sulfinol its advantages.
Sulfinol with DIPA (Sulfinol-D) is used when complete removal of H2S, CO2, and COS is desired. Sulfinol with MDEA (Sulfinol-M) is used for the selective removal of H2S in the presence of CO2, with partial removal of COS. Both Sulfinols can reduce the total sulfur content of treated gas down to low ppm levels. Some disadvantages are: a higher heavy hydrocarbon co-absorption, and a reclaimer is sometimes required for Sulfinol D when removing CO2.
The presence of the physical solvent, Sulfolane@, allows higher acid gas loadings compared to systems based on amine only. Typical loadings are 1.5 mol of acid gas per mole of Sulfinol solution. Higher acid gas loadings, together with a lower energy of regeneration, can result in lower capital and energy costs per unit of acid gas removed as compared to the ethanolamine processes.
Features of the Sulfinol process include essentially complete removal of mercaptans, high removal rate of COS, lower foaming tendency, lower corrosion rates, and the ability to slip up to 50% CO2.
The design is similar to that of the ethanolamines. Degradation of DIPA to oxazolidones (DIPA-OX) usually necessitates the installation of a reclaimer for their removal.
As with the ethanolamine processes, aromatics and heavy hydrocarbons in the feed gas should be removed prior to contact with the Sulfinol solution to minimize foaming.
The merits of the Sulfinol process, as compared to the ethanolamine processes, are many, but there are other factors that must be considered before selecting the appropriate gas treating process. For example, a licensing fee, while not necessary for the ethanolamine processes, is required for the Sulfinol process. The solvent costs are generally higher for the Sulfinol process than they are for DEA. Operators are more familiar with DEA and the typical problems associated with this process. In cases of low acid gas partial pressure, the advantage of a lower circulation rate for the Sulfinol process diminishes compared to the DEA process.
6.5.2 Selefining Process
This process developed by Snamprogetti uses tertiary amines such as dimethyl ethanol amine (DMEA) dissolved in an organic solvent with limited amounts of water. The process is very selective to H2S as CO2 hydration is almost completely avoided. It also removes other sulfur species such as mercaptans, COS and CS2. It has a tendency to co-absorb hydrocarbons which can to some extent be controlled by increasing the water content of the solvent.
6.3.1.1 Process Description
A typical amine system is shown in Figure 6.3. Sour gas enters the system through an inlet separator (scrubber) to remove any entrained water or hydrocarbon liquids. Gas enters the bottom of the amine absorber (contactor) and flows upward, countercurrent to the lean amine solution which flows down from the top. The lean amine that returns to the contactor is maintained at a temperature above the vapor that exits the contactor to prevent any condensation of heavier liquid hydrocarbons. Intimate contact between the gas and amine solution is achieved by use of either trays or packing in the contactor.
The absorber tower consists of trays (diameters >20 in. (500 mm)), conventional packing (diameters <20 in.) or structured packing (diameters >20 in.).
Sweetened gas leaves the top of the tower. An optional outlet (separator) scrubber may be included to recover entrained amine from the sweet gas. Since the natural gas leaving the top of the tower is saturated with water, the gas will require dehydration before entering a pipeline.
Rich amine, solution containing CO2 and H2S, leaves the bottom of the absorber and flows to the flash tank (drum) where most of the dissolved hydrocarbon gases or entrained hydrocarbon condensates are removed. A small amount of the acid gases flash to the vapor phase. From the flash drum, the rich amine proceeds to the rich amine/lean amine heat exchanger where it recovers some of the sensible heat from the lean amine stream, which decreases the heat duty on the amine reboiler and the solvent cooler. Preheated rich amine then enters the amine stripping tower where heat from the reboiler breaks the bonds between the amine and acid gases. Acid gases are removed overhead and lean amine is removed from the bottom of the stripper.
Hot lean amine flows to the rich amine/lean amine heat exchanger and then to additional coolers, typically aerial coolers, to lower its temperature about 100F (5.50C) above the inlet gas temperature. This reduces the amount of hydrocarbons condensed in the amine solution when the amine contacts the sour gas.
A side stream of amine, of about 3%, is taken off after the rich/lean amine heat exchanger and is flowed through a charcoal filter to clean the solution of contaminants (not included in figure). Cooled lean amine is then pumped up to the absorber pressure and enters the top of the absorber. Amine solution flows down the absorber where it absorbs the acid gases. Rich amine is then removed at the bottom of the tower and the cycle is repeated.
6.3.1.2 General Remarks
The contactor operates above ambient temperature because of the combined exothermic heat of absorption and reaction. The maximum temperature is in the lower portion of the tower because the majority of the absorption and reaction occurs near the bottom of the unit. The temperature “bulge” in the tower can be up to about 180°F (80°C). The treated gas leaves the top of the tower water saturated and at a temperature controlled by the temperature of the lean amine that enters, usually around 100°F (38°C).
The sweet gas leaving the contactor is saturated with water so dehydration, is normally required prior to sale. If MEA is the sweetening agent, or the contactor is operating at unusually high temperature, a water wash may be used to attempt to recover some of the vaporized and/or entrained amine from the gas leaving the contactor. If a water wash is used it generally will consist of three or four trays at the top of the contactor, with makeup water to the unit being used as the wash liquid.
Acid gas stripped from the amine passes out of the top of the stripper. It goes through a condenser and separator to cool the stream and recover water. The recovered water is usually returned to the stripper as reflux. The acid gas from the reflux separator is either vented, incinerated, sent to sulfur recovery facilities, or compressed for sale or reinjected into a suitable reservoir for enhanced oil recovery projects or for sequestration.
Fig. 6-3. Gas Sweetening by Amine absorption, operating conditions are representative, not definitive.
6.3.1.3 Operating Issues
Corrosion—Some of the major factors that affect corrosion are:
• Amine concentration (higher concentrations favor corrosion)
• Rich amine acid gas loading (higher gas loadings in the amine favor corrosion)
• Oxygen concentration
• Heat stable salts (higher concentrations promote corrosion and foaming)
In addition to destroying vessels and piping, the corrosion products can cause foaming.
Solution Foaming—Foaming of the liquid amine solution is a major problem because it results in poor vapor−liquid contact, poor solution distribution, and solution holdup with resulting carryover and off spec gas. Among the causes of foaming are suspended solids, liquid hydrocarbons, surface active agents, such as those contained in inhibitors and compressor oils, and amine degradation products, including heat stable salts. One obvious cure is to remove the offending materials; the other is to add antifoaming agents.
Heat Stable Salts (HSS) —As mentioned above, these amine degradation products can cause both corrosion and foaming. They are normally dealt with through the use of amine reclaimers.
6.3.1.4 Reclaimer
A reclaimer is usually required for MEA and DGA® amine-based systems. The reclaimer helps remove degradation products from the solution and also aids in the removal of heat stable salts, suspended solids, acids and iron compounds. The reclaimers in MEA and DGA® systems differ. For MEA, a basic solution helps reverse the reactions. Soda ash and/or caustic soda is added to the MEA reclaimer to provide a pH of approximately 8-9; no addition is required for the DGA® reclaimer system. Reclaimers generally operate on a side stream of 1-3% of the total amine circulation rate. Reclaimer sizing depends on the total inventory of the plant and the rate of degradation expected.
Reclaimer operation is a semi-continuous batch operation. The reclaimer is filled with hot amine solution and, if necessary, soda ash is added. As the temperature in the reclaimer increases, the liquid will begin to distill. Overhead vapors can be condensed and pumped back into the amine system, but generally the reclaimer is operated at slightly above stripper column pressure and the vapors are returned to the stripper.
The initial vapor composition is essentially water. Continued distillation will cause the solution to become more and more concentrated with amine. This raises the boiling point of the solution and amine will begin to distill overhead. Fresh feed is continually added until the boiling point of the material in the reclaimer reboiler reaches 280° to 300°F. At this point, distillation is continued for a short time adding only water to help recover residual amine in the reclaimer reboiler. The reclaimer is then cleaned, recharged, and the cycle is repeated.
Reclaimer "sludge" removed during cleaning must be handled with care. Disposal of the "sludge" must be in accordance with the governing regulations. If needed a reclaiming company may be contracted to remove degradation products or heat stable salts from the amine.
DEA does not form a significant amount of nonregenerable degradation products, and it requires more difficult reclaiming through vacuum distillation or ion exchange.
6.3.1.5 Basic Amine Chemistry
Amines are bases, and the important reaction in gas processing is the ability of the amine to form salts with the weak acids formed by H2S and CO2 in an aqueous solution. When a gas stream that contains the H2S, CO2, or both, is contacted by a primary or secondary amine solution, the acid gases react to form a soluble acid−base complex, a salt, in the treating solution. The reaction between the amine and both H2S and CO2 is highly exothermic. Regardless of the structure of the amine, the overall reactions between H2S and amines are simple since H2S reacts directly and rapidly with all amines to form the bisulfide and the sulfide by Eqs. 6-1 & 6-2:
For hydrogen sulfide removal
RNH2 + H2S ↔ RNH3+ + HS– Fast Eq. 6-1
RNH2 + HS– ↔ RNH3+ + S–2 Fast Eq. 6-2
For carbon dioxide removal
2RNH2 + CO2 ↔ RNH3+ + RNHCOO– Fast Eq. 6-3
RNH2 + CO2 + H2O ↔ RNH3+ + HCO3– Slow Eq. 6-4
RNH2 + HCO3– ↔ RNH3+ + CO3–2 Slow Eq. 6-5
Concerning the chemical reactions with CO2, primary amines (RNH2) such as MEA and DGA® agent, and secondary amines (RR’NH) such as DEA and DIPA, differ from tertiary amines (RR’R"N) such as TEA and MDEA.
For the reactions discussed above, high pressures and low temperatures drive the reactions to the right, whereas high temperatures and low pressures favor the reverse reaction, which thus provides a mechanism for regeneration of the amine solution.
Primary and Secondary Amines
With the primary and secondary amines, the predominant overall reaction (Eq. 6-3) rapidly leads to the formation of a stable carbamate which is slow to further hydrolyze to bicarbonate.
The other overall reactions leading to bicarbonate (Eq. 6-4) and to carbonate (Eq. 6-5) are slow because they have to proceed through the hydration of CO2.
Therefore, according to Eq. 6-3 there is a theoretical limit to the chemical loading capacity of the primary and secondary amine solutions to 0.5 mole CO2 per mole of amine, even at relatively high partial pressures of CO2 in the gas to be treated.
For primary and secondary amines, little difference exists between the H2S and CO2 reaction rates because of the availability of the rapid carbamate formation for CO2 absorption. Therefore, the primary and secondary amines achieve essentially complete removal of H2S and CO2.
Tertiary Amines
Unlike primary and secondary amines, the nitrogen (N) in tertiary amines (RR’R"N) has no free hydrogen ( H ) to rapidly form carbamate as per overall Eq. 6-3. Consequently, the removal of CO2 by tertiary amines can only follow the slow route to bicarbonate by Eq. 6-4 and carbonate by Eq. 6-5.
The slowness of the reaction leading to bicarbonate is the underlying reason why tertiary amines can be considered selective for H2S removal, by playing with absorption contact time, and this attribute can be used to full advantage when complete CO2 removal is not necessary.
However, the slow route to bicarbonates theoretically allows at equilibrium a chemical loading ratio of one mole of CO2 per mole of amine. Furthermore, at high partial pressure, the solubility of CO2 in tertiary amines is far greater than in the primary and secondary amines thus further enhancing the CO2 loading by physical solubility at high partial pressure.
Therefore, in case of gases to be treated for bulk CO2 removal, large amounts of CO2 can be liberated from the rich solvent by simple flash alleviating the thermal regeneration duty with consequent energy savings. In other words, because the tertiary amines have no labile hydrogen, they cannot form the carbamate. Tertiary amines must react with CO2 via the slow hydrolysis mechanism in Equation 6.4. With only the slow acid−base reaction available for CO2 absorption, MDEA (methyldiethanolamine) and several of the formulated MDEA products yield significant selectivity toward H2S relative to CO2, and, consequently, all of the H2S is removed while some of the CO2 “slips” through with the gas.
Activated Tertiary Amines
The use of activators mitigates the slowness of the reaction to bicarbonate for tertiary amines. Activators are generally primary or secondary amines; they are tailored to increase both the hydrolysis of the carbamate and the rate of hydration of dissolved CO2 thus making the activated-tertiary amines specially suitable for efficient and economical bulk CO2 removal when selectivity is not required (see section on MDEA).
6.3.1.6 Amines Used
Monoethanolamine
MEA is a primary amine, which has had widespread use as a gas sweetening agent. The process is well proven and can meet pipeline specifications. MEA is a stable compound and, in the absence of other chemicals, suffers no degradation or decomposition at temperatures up to its normal boiling point.
Gas sweetening with monoethanolamine (MEA) is used where there are low contactor pressures and/or stringent acid gas specifications. MEA removes both H2S and CO2 from gas streams. H2S concentrations well below 4.0 ppmv can be achieved. CO2 concentrations as low as 100 ppmv can be obtained at low to moderate pressures.
COS and CS2 are removed by MEA, but the reactions are irreversible unless a reclaimer is used. Even with a reclaimer, complete reversal of the reactions may not be achieved. The result is solution loss and build-up of degradation products in the system. Total acid gas pick up is traditionally limited to 0.3-0.35 moles of acid gas/mole of MEA and solution concentration is usually limited to 10-20 wt%. Inhibitors can be used to allow much higher solution strengths and acid gas loadings. Because MEA has the highest vapor pressure of the amines used for gas treating, solution losses through vaporization from the contactor and stripper can be high (MEA losses of 1-3 lbs/MMscf (16-48 kg/ MM m3) of inlet gas). This problem can be minimized by using a water wash.
MEA reactions are reversible by changing the system temperature. Reactions with CO2 and H2S are reversed in the stripping column by heating the rich MEA to about 2450F at 10 psig (118 0C at 69 kPa). Acid gases evolve into the vapor and are removed from the still overhead. Thus, the MEA is regenerated.
MEA systems foam rather easily resulting in excessive amine carryover from the absorber. Foaming can be caused by a number of foreign materials such as condensed hydrocarbons, degradation products, solids such as carbon or iron sulfide, excess corrosion inhibitor, and valve grease.
A microfiber filter separator should be installed at the gas inlet to the MEA contactor. It is an effective method of foam control and removes many of the contaminants before they enter the system. Hydrocarbon liquids are usually removed in the flash tank. Degradation products are removed in a reclaimer as described above.
A gas blanket system is installed on MEA storage tanks and surge vessels. This prevents oxidation of MEA. Sweet natural gas or nitrogen is normally used.
Disadvantages include the reaction of MEA with carbonyl sulfide (COS) and carbon disulfide (CS2) to form heat-stable salts, which cannot be regenerated at normal stripping column temperatures. At temperatures above 245 0F (118 0C) a side reaction with CO2 exists that produces oxazolidone-2, a heat-stable salt, which consumes MEA from the process. Normal regeneration temperature in the still will not regenerate heat-stable salts or oxazolidone-2. A reclaimer is often included to remove these contaminants.
Diethanolamine
This process employs an aqueous solution of diethanolamine (DEA) (a secondary amine). DEA will not treat to pipeline quality gas specifications at as low a pressure as will MEA.
This process is used for high pressure, high acid gas content streams having a relatively high ratio of H2S/CO2.
Although mole/mole loadings as high as 0.8-0.9 have been reported, most conventional DEA plants still operate at significantly lower loadings.
The process flow scheme for conventional DEA plants resembles the MEA process. The advantages and disadvantages of DEA as compared to MEA are:
• The mole/mole loadings typically used with DEA (0.35-0.82 mole/mole) are much higher than those normally used (0.3-0.4) for MEA.
• Molecular weight of DEA is 105 compared to 61 for MEA. Combination of molecular weights and reaction stoichiometry means that about 1.7 lbs (0.77 kg) of DEA must be circulated to react with the same amount of acid gas as 1.0 lbs (0.45 kg) of MEA. The solution strength of DEA ranges up to 35% by weight compared to 20% for MEA. Loadings for DEA systems range from 0.35 to 0.65 mol of acid gas per mole of DEA without excessive corrosion. The result of this is that the circulation rate of a DEA solution is slightly less than in a comparable MEA system.
• Because DEA does not form a significant amount of nonregenerable degradation products, a reclaimer is not required.
• DEA is a secondary amine and is chemically weaker than MEA, and less heat is required to strip the amine solution.
• DEA forms a regenerable compound with COS and CS2 and can be used for the partial removal of COS and CS2 without significant solution losses.
• As a secondary amine, DEA is less alkaline than MEA. DEA systems do suffer the same corrosion problems, but not as severely as those using MEA. Solution strengths are typically from 25% to 35% DEA by weight in water.
• Vapor pressure of DEA is about 1/30 of the vapor pressure of MEA. Thus, DEA amine losses are much lower than in an MEA system.
Diglycolamine
DGA® is a primary amine capable of removing not only H2S and CO2, but also COS and mercaptans from gas and liquid streams. DGA® has been used to treat natural gas to 4.0 ppmv at pressures as low as 125 psig. Compared with MEA, low vapor pressure allows Diglycolamine [ 2-(2-aminoethoxy) ethanol] (DGA) is reclaimed on-site to remove heat stable salts and reaction products with COS and CS2.
DGA® has a greater affinity for the absorption of aromatics, olefins, and heavy hydrocarbons than the MEA and DEA systems.
Therefore, adequate carbon filtration should be included in the design of a DGA® treating unit.
The process flow for the DGA® treating process is similar to that of the MEA treating process. The major differences are:
• Higher acid gas pick-up per gallon of amine can be obtained by using 50-60% solution strength rather than 15-20% for MEA (more moles of amine per volume of solution).
• The required treating circulation rate is lower. This is a direct function of higher amine concentration.
• Reduced reboiler steam consumption.
• DGA® has an advantage for plants operating in cold climates where freezing of the solution could occur. The freezing point for 50% DGA® solution is –30°F.
• Low vapor pressure decreases amine losses.
• Unlike MEA, degradation products from reactions with COS and CS2 can be regenerated.
Methyldiethanolamine
Methyldiethanolamine (MDEA) is a tertiary amine which can be used to selectively remove
H2S to pipeline specifications at moderate to high pressure. If increased concentration of CO2 in the residue gas does cause a problem with contract specifications or downstream processing, further treatment will be required.
If the gas is contacted at pressures ranging from 800 to 1000 psig, H2S levels can be reduced to concentrations required by pipelines. While at the same time, 40-60% of the CO2 present flows through the contactor, untreated.
The H2S/CO2 ratio in the acid gas can be 10-15 times as great as the H2S/CO2 ratio in the sour gas. Some of the benefits of selective removal of H2S include:
• Reduced solution flow rates resulting from a reduction in the amount of acid gas removed.
• Smaller amine regeneration unit. Significant capital savings are realized due to reduced pump and regeneration requirements.
• Higher H2S concentrations in the acid gas resulting in reduced problems in sulfur recovery.
• MDEA has a lower heat requirement due to its low heat of regeneration. In some applications, energy requirements for gas treating can be reduced as much as 75% by changing from DEA to MDEA.
• It is not reclaimable by conventional methods.
CO2 hydrolyzes much slower than H2S. This makes possible significant selectivity of tertiary amines for H2S. This fact is used for selective removal of H2S from gases containing both H2S and CO2.
A feature of MDEA is that it can be partially regenerated in a simple flash. As a consequence the removal of bulk H2S and CO2 may be achieved with a modest heat input for regeneration.
However, as MDEA solutions react only slowly with CO2, activators must be added to the MDEA solution to enhance CO2 absorption and the solvent is then called activated MDEA.
Solution strengths typically range from 40% to 50% MDEA by weight. Acid gas loading varies from 0.2 to 0.4 or more moles of acid gas per mole of MDEA depending on the supplier.
Diisopropanolamine
Diisopropanolamine (DIPA) is a secondary amine which exhibits, though not as great as tertiary amines, selectivity for H2S. It is similar to DEA systems but offers the following advantages:
• Carbonyl sulfide (COS) can be removed and the DIPA solution can be regenerated easily
• The system is generally noncorrosive and has a lower energy consumption.
• At low pressures, DIPA will preferentially remove H2S. As pressure increases, the selectivity of the process decreases and DIPA removes increasing amounts of CO2. Thus, this system can be used either to selectively remove H2S or to remove both CO2 and H2S.
Formulated Solvents and Mixed Amines
The selectivity of MDEA can be reduced by addition of MEA, DEA, or proprietary additives. Thus, it can be tailored to meet the desired amount of CO2 slippage and still have lower energy requirements than do primary and secondary amines.
Formulated Solvents is the name given to a new family of amine based solvents. Their popularity is primarily due to equipment size reduction and energy savings over most of the other amines. All the advantages of MDEA are valid for the Formulated Solvents, usually to a greater degree. Some formulations are capable of slipping larger portions of inlet CO2 (than MDEA) to the outlet gas and at the same time removing H2S to less than 4 ppmv. For example, under conditions of low absorber pressure and high CO2 /H2S ratios, such as Claus tail gas clean-up units, certain solvent formulations can slip upwards to 90 percent of the incoming CO2 to the incinerator.
While at the other extreme, certain formulations remove CO2 to a level suitable for cryogenic plant feed. Formulations are also available for CO2 removal in ammonia plants. Finally, there are solvent formulations which produce H2S to 4 ppmv pipeline specifications, while reducing high inlet CO2 concentrations to 2% for delivery to a pipeline. This case is sometimes referred to as bulk CO2 removal.
This need for a wide performance spectrum has led “Formulated Solvent” suppliers to develop a large stable of different MDEA-based solvent formulations. Most Formulated Solvents are enhancements to MDEA discussed above. Thus, they are referred to as MDEA-based solvents or formulations.
Benefits claimed by suppliers are:
For New Plants
• reduced corrosion
• reduced circulation rate
• lower energy requirements
• smaller equipment due to reduced circulation rates
For Existing Plants
• increase in capacity, i.e., gas through-put or higher inlet acid gas composition
• reduced corrosion
• lower energy requirements and reduced circulation rates
Inhibited Amine Systems
Inhibited amine processes use standard amines that have been combined with special inhibiting agents that minimize corrosion. They allow higher solution concentrations and higher acid gas loadings, thus reducing required circulation rates and energy requirements. They also utilize hot potassium carbonate to remove CO2 and H2S. As a general rule, this process should be considered when the partial pressure of the acid gas is 20 psia (138 kPa) or greater. The process is not recommended for low-pressure absorption, or high-pressure absorption of low-concentration acid gas.
Sterically Hindered Amines
In acid gas removal, steric hindrance involves alteration of the reactivity of a primary or secondary amine by a change in the alkanol structure of the amine.
A large hydrocarbon group attached to the nitrogen shields the nitrogen atom and hinders (inhibits) the carbamate reaction. The H2S reaction is not significantly affected by amine structure, because the proton is small and can reach the nitrogen. However, CO2 removal can be significantly affected if the amine structure hinders the fast carbamate formation reaction and allows only the much slower bicarbonate formation.
Table 6.7. Some Representative Operating Parameters for Amine Systems
6.3.1.7 Heats of Reaction
Amines are basic solutions. These solutions react with the hydrogen sulfide and carbon dioxide to form a salt. The process of absorbing the acid gases generates heat.
The magnitude of the exothermic heats of reaction, which includes the heat of solution, of the amines with the acid gases is important because the heat liberated in the reaction must be added back in the regeneration step. Thus, a low heat of reaction translates into smaller energy regeneration requirements. Table 6.8 summarizes the important data for the common amines. The values in Table 6.8 are approximate because heats of reaction vary with acid gas loading and solution concentration.
Table 6.8. Average Heats of Reaction of the Acid Gases in Amine Solutions
6.3.2 Alkaline Salt Process (Hot Carbonate)
The hot potassium carbonate process for removing CO2 and H2S was developed by the United States Bureau of Mines and is described by Benson and coworkers in two papers. Although the process was developed for the removal of CO2, it can also remove H2S if H2S is present with CO2 (not suitable for gas streams containing only H2S.). Special designs are required for removing H2S to pipeline specifications or to reduce CO2 to low levels.
The process is very similar in concept to the amine process, in that after physical absorption into the liquid, the CO2 and H2S react chemically with the solution. The chemistry is relatively complex, but the overall reactions are represented by
K2CO3 + CO2 + H2O ↔ 2KHCO3 Eq. 6-6
K2CO3 + H2S ↔ KHS + KHCO3 Eq. 6-7
In a typical application, the contactor will operate at approximately 300 psig (20 barg), with the lean carbonate solution entering near 225°F (110°C) and leaving at 240°F (115°C). The rich carbonate pressure is reduced to approximately 5 psig (0.3 barg) as it enters the stripper.
Approximately one third to two thirds of the absorbed CO2 is released by the pressure reduction, reducing the amount of steam required for stripping. The lean carbonate solution leaves the stripper at the same temperature as it enters the contactor, and eliminates the need for heat exchange between the rich and lean streams. The heat of solution for absorption of CO2 in potassium carbonate is small, approximately 32 Btu/cu ft of CO2, and consequently the temperature rise in the contactor is small and less energy is required for regeneration.
Potassium carbonate processes are somewhat effective in removing carbonyl sulfide and carbon disulfide. Potassium carbonate works best on gas streams with a CO2 partial pressure of 30-90 psi.
Pipeline-quality gas often requires secondary treating using an amine or similar system to reduce H2S level to 4 ppm.
Because this system is operated at high temperatures to increase the solubility of carbonates, the designer must be careful to avoid dead spots in the system where the solution could cool and precipitate solids. If solids do precipitate, the system may suffer from plugging, erosion, or foaming.
Hot potassium carbonate solutions are corrosive. All carbon steel must be stress relieved to limit corrosion. Varieties of corrosion inhibitors, such as fatty amines or potassium dichromate, are available to decrease corrosion rates. There are three basic process flow variations for the potassium carbonate process. The flow scheme required depends on the outlet specification of the natural gas. These are:
Single Stage Process
The single stage process is shown in Fig. 6-4. Potassium carbonate is pumped to the top of a packed or trayed contactor where it contacts the gas stream. The rich solution flows to the stripper where the acid gases are stripped with steam. The lean solution is then pumped back to the contactor to complete the cycle.
Split Flow Process
In this process scheme (Fig. 6-5) the lean solution stream is split. Hot solution is fed to the middle of the contactor for bulk removal. The remainder is cooled to improve equilibrium and is fed to the top of the contactor for trim acid gas removal.
Two Stage Process
In this process scheme (Fig. 6-6) the contactor is like that of the split flow process. In addition, the stripper is in two sections. A major portion of the solution is removed at the midpoint of the stripper and pumped to the lower section of the contactor.
Numerous improvements have been made to the potassium carbonate process resulting in significant reduction in capital and operating costs. At the same time, lower acid gas concentration in the treated gas can now be achieved. The most popular of the carbonate processes are:
Benfield Process
The Benfield Process is licensed by UOP. Several activators are used to enhance the performance of the potassium carbonate solution.
Hi-Pure Process
The Hi-Pure process is a combination conventional Benfield potassium carbonate process and alkanolamine process. The gas stream is first contacted with potassium carbonate followed by contacting with an amine. The process can achieve outlet CO2 concentrations as low as 30 ppmv and H2S concentrations of 1 ppmv.
Catacarb Process
The Catacarb Process is licensed by Eickmeyer and Associates. Activators, corrosion inhibitors, potassium salts, and water are contained in the solution. This process is mostly used in the ammonia industry.
Fig. 6-4. Alkaline Salt: Single-Stage Process
Fig. 6-5. Alkaline Salt: Split-Flow Process
Fig. 6-6. Alkaline Salt: Two-Stage Process
6.3.3 Specialty Batch Chemical Solvents
Several batch chemical processes have been developed and have specific areas of application.
Processes include
General Process Description
Gas is flowed into a vessel and contacted with the solvent. Acid components are converted to soluble salts, which are nonregenerable, limiting the life of the solution.
Once saturation levels are reached, the solution must be replaced.
For some of these processes, the spent solutions are not hazardous, but for others, the spent solutions have been labeled hazardous and, if used, must be disposed of as Class IV materials.
Units in these processes have a wide operating range, with acid gas concentrations ranging from as low as 10 ppm to as high as 20%. Operating pressures range from near atmospheric to >1000 psig (7000 kPa). Some units have been designed to handle from several thousand cubic feet per day to more than 15 MMscfd (several hundred cubic meters per day to more than 420,000 m3 per day).
a- Sulfa-Check
Sulfa-Check is a product of ExxonMobil Chemicals which selectively removes H2S and mercaptans from natural gas in the presence of CO2. The patented process converts the sour gas directly to sulfur. This is accomplished by sparging “bubbling” the gas in a buffered, water based oxidizing solution containing sodium nitrite (NaNO2). The sodium nitrite is reduced to ammonia (NH3) which remains in solution. The spent product is classified as non-hazardous.
Dozens of field applications include gas rates ranging from 15 Mscf/d to 3.0 MMscf/d and inlet H2S concentrations ranging from 10 to 3000 ppmv.
Reaction rate is independent of the concentration of the oxidizing agent.
There is no limit to the concentration of H2S treated. Process is most economical for acid gas streams containing from1 ppm to 1% H2S. pH must be held above 7.5 to control selectivity and optimize H2S removal. One gallon (4 l) of oxidizing solution can remove up to 2 lbs (1 kg) of H2S when the system is operated at ambient temperatures<100 0F (38 0C). If gas temperatures exceed 100 0F (38 0C), the solubility of sulfur in the oxidizing agent decreases.
Operating pressure of at least 20 psig is required for proper unit operation to maintain bubble flow through the column. Bubble flow is necessary to produce intimate mixing of the gas and liquid.
Oxidizing solution will eventually become saturated and require replacement. Disposal of this slurry poses no environmental problem, as the reaction produces an aqueous slurry of sulfur and sodium salt.
A number of variables, including some associated risks must be considered prior to determining if the Sulfa-Check process is applicable. For example, low levels of ammonia may appear in the treated gas. Also, the reduction of NO2 may result in the formation of nitric oxide (NO). If air is present in the raw gas, it will react with nitric oxide to form nitrogen dioxide (NO2). NO2 is a strong oxidizing agent that will react with elastomers and odorants and cause corrosion in a moist environment.
b- Caustic Wash
Caustic (NaOH) scrubbing systems can be used to treat natural gas streams to remove CO2, CS2, H2S, and mercaptans. The process employs countercurrent contacting of the gas stream with a caustic solution in a packed or trayed column. The column may contain one stage or several stages depending on the required degree of removal. The multi-stage systems generally have different caustic concentrations ranging from 4-6 weight percent in the first stage to 8-10 weight percent in the latter stages. Multiple stages increase the caustic efficiency while maintaining a sufficient driving force to achieve absorption.
The spent solution is either regenerated or discarded depending on what acid gas components are present in the gas stream. If only mercaptans are present, the caustic solution is regenerated with steam in a stripping still. If CO2 is present, a nonregenerable product (Na2CO3) is formed and the solution must be discarded. As a result, the presence of CO2 in caustic systems leads to high caustic consumption. This is a serious disadvantage of the caustic scrubbing process. The spent caustic solutions are considered hazardous wastes. Natural gas is usually water washed after a caustic wash to remove any caustic entrained in the gas prior to dehydration.
The chemical reactions involved are as follows:
H2S + 2 NaOH → Na2S + 2H2O Eq 6-8
RSH + NaOH → RSNa + H2O Eq 6-9
CO2 + 2 NaOH → Na2CO3 + H2O Eq 6-10
CS2 + 2 NaOH → 2 NaHS + CO2 Eq 6-11
Fig. 6-7. Regenerative Caustic
c- Sulfa-Scrub
Sulfa-Scrub from Baker-Petrolite is a product that uses triazine compound to selectively react with H2S. The raw gas can be sparged into a vessel containing the liquid scavenging agent. Alternatively the liquid may be injected to the gas stream as an H2S scavenger. The spent material is considered as non-hazardous and is claimed to be an excellent corrosion inhibitor.
6.4 Physical Solvent Processes
In the amine and alkali salt processes, the acid gases are removed in two steps:
physical absorption followed by chemical reaction. In processes such as Selexol® or Rectisol®, no chemical reaction occurs and acid gas removal depends entirely on physical absorption. Some of the inherent advantages and disadvantages of physical absorption processes are summarized below:
These processes flow scheme is as shown in Fig. 6-8. In general, a physical solvent process should be considered when:
1- The partial pressure of the acid gas in the feed is greater than 50 psi.
2- The heavy hydrocarbon concentration in the feed gas is low.
3- Bulk removal of the acid gas is desired.
4- Selective removal of H2S is desired.
These processes are economically attractive because little energy is required for regeneration. The solvents are regenerated by:
1- Multi-stage flashing to low pressures.
2- Regeneration at low temperatures with an inert stripping gas.
3- Heating and stripping of solution with steam/solvent vapors.
In general, physical solvents are capable of removing COS, CS2, and mercaptans.
In certain instances, physical absorption processes are capable of simultaneously dehydrating and treating the gas although additional equipment and higher energy requirements may be needed to dry the solvent. The processes operate at ambient or subambient temperature to enhance the solubility of the acid gases. The solvents are relatively noncorrosive so carbon steel can be used. Chemical losses are low due to low solvent vapor pressure or refrigerated conditions. Physical solvents will absorb heavy hydrocarbons from the gas stream resulting in high hydrocarbon content in the acid gas stream as well as possibly significant hydrocarbon losses. Some of the physical absorption processes are summarized below.
Fig. 6-8. Typical Gas Sweetening by Physical Absorption
6.4.1 Selexol®
This process developed by Allied Chemical Corp. uses a polyethylene glycol derivative as a solvent. The solvent is selective for RSH, CS2, H2S, and other sulfur compounds. The process can be used to selectively or simultaneously remove sulfur compounds, carbon dioxide, water, as well as paraffinic, olefinic, aromatic and chlorinated hydrocarbons from a gas or air stream. Because water and heavy hydrocarbons are highly soluble in Selexol®, the treated gas from a Selexol® unit normally meets both water and hydrocarbon dew point specifications.
Levels of CO2 can be reduced by about 85%. The process is economical when high acid gas partial pressures exist and there is an absence of heavy ends in the gas.
This process will not typically remove enough CO2 to meet pipeline gas requirements. DIPA can be added to the solution to remove CO2 down to pipeline specifications. The process also removes water to <7 lbs/MMscf (0.11 g/std m3). The addition of DIPA increases the relatively low stripper heat duty.
The vendor states that the solvent is very stable, no degradation products are formed or disposed of, and no solvent reclaiming is required. Depending on the applications, the operating pressure could be as low as ambient though higher pressure is preferred. Operating temperature varies from 0°F to ambient. Selexol® is a Dow Chemical solvent and a UOP technology.
Table 6-9 presents RK, the ratio of the K-value* for methane, KCH4 (arbitrarily assigned a value of 1), to the K-values of the other component, RK = KCH4/Kcomponent . (The K-value is the ratio of the mole fraction of the component in the vapor phase (y) to its mole fraction in the liquid phase (x), K = y/x. High K-values indicate the material is predominately in the vapor phase, whereas low K-values indicate a higher concentration in the liquid phase (x)).
Table 6-9 Typical Relative Ratio of K-values
An R value greater than1 indicates that the solubility of the component in Selexol® is greater than that of methane. The values should be regarded as only representative because pressure and temperature are not specified and, as previously noted, the composition of Selexol® is variable.
Because RK for CO2 and H2S are 15 and 134, respectively, these gases are preferentially absorbed (relative to CH4), and, consequently, physical absorption is an effective technique for acid gas removal. The process can reduce H2S to 4 ppmv, reduce CO2 to levels below 50 ppmv, and essentially remove all mercaptans, CS2, and COS. Two additional features of Table 6-9 are worth mentioning. Because the RK values for hydrocarbons heavier than CH4 are fairly high (6.4 for C2H6, 15.3 for C3H8, and 35 for n-C4H10), Selexol® will remove substantial quantities of these hydrocarbons, a feature that can be either positive or negative, depending on the composition of the gas being processed and the desired products. Finally, the RK value of H2O is extremely high and consequently, Selexol® provides some dehydration.
Volumes of some gases in scf absorbed/gal Selexol® are plotted in Figure 6-9. The figure assumes a Henry’s law relationship, (For an ideal system, Henry’s law assumes a linear relation between the solubility of gas component (i) and its partial pressure, yiP = kixi where ki is the Henry’s constant.) which provides approximate solubility at higher pressures. The lines also ignore probable Interaction between solutes.
Applications of Selexol® are varied and, consequently, no common process flow diagrams are available. One plant is shown in Figure 6-10. The plant pretreats the gas to reduce CO2, ethane, and heavier hydrocarbon levels before final purification in molecular sieve units and subsequent liquefaction. The plant is designed to process 26 MMscfd (0.74 MMSm3/d) entering the Selexol® unit at 603 psia (41.6 bar) and 32°F (0°C). The lean solvent, cooled to 25°F (−3.9°C) with propane refrigerant, enters the absorber where it absorbs CO2 and some of the ethane and heavier hydrocarbons. The rich solvent from the absorber is regenerated by reduction of the pressure in three flash drums, from 603 to 106 psia (41.6 to 7.3 bar) in the high-pressure drum, from 106 to 16 psia (7.3 to 1.1 bar) in the medium-pressure drum, and from 16 to 3 psia (1.1 to 0.21 bar) in the vacuum drum. Lean Selexol® from the vacuum drum is recompressed and sent to the propane chiller. The treated gas that leaves the absorber passes through a knockout drum and filter separator to remove entrained Selexol® and condensed hydrocarbons. Table 6-11 shows that the treated gas meets the specifications of a maximum of 0.50% CO2 and a maximum of 6.5% ethane and heavier hydrocarbons. In addition, the water content of the gas is reduced from 75 ppmv to 12 ppmv, H2S is reduced from 2 ppmv to essentially nothing, and methyl mercaptan is reduced from 5 ppmv to 1 ppmv. Unlike the amine systems, no irreversible products are generated in the process, which thus eliminates the need for reclaiming.
Table 6-10. Representative Property Data for Selexol®
Fig. 6-9. Solubility of various gases in Selexol® solvent at 70°F (21°C) as a function of partial pressure.
Table 6-11. Composition of Inlet and Outlet Gas in a Selexol® Unit
Fig. 6-10. Process schematic for a Selexol® gas treating facility.
6.4.2 Fluor Solvent
The Fluor Solvent process uses propylene carbonate as a physical solvent to remove CO2 and H2S. Propylene carbonate also removes C3+ hydrocarbons, COS, SO2, CS2, and H2O from the natural gas stream. Thus, in one step the natural gas can be sweetened and dehydrated to pipeline quality.
This process is used for bulk removal of CO2 and is not used to treat to <3% CO2.
This system requires special design features such as larger absorbers and higher circulation rates to obtain pipeline quality and usually is not economically applicable for outlet requirements.
Propylene carbonate has the following characteristics, which make it suitable as a solvent for acid gas treating:
• High degree of solubility for CO2 and other gases
• Low heat of solution for CO2
• Low vapor pressure at operating temperature
• Low solubility for light hydrocarbons (C1, C2)
• Chemically nonreactive toward all natural gas components
• Low viscosity
• Noncorrosive toward common metals
The above characteristics combine to yield a system that has low heat and pumping requirements, is relatively noncorrosive, and suffers only minimal solvent losses (<1 lbs/MMscf).
Solvent temperatures below ambient are often used to increase solvent gas capacity and, therefore, decrease circulation rates. Expansion of the rich solvent and flash gases through power turbines can provide the required refrigeration. Alternately, auxiliary refrigeration may be included to further decrease circulation rates.
6.4.3 Rectisol Process®
This process uses pure refrigerated methanol as a solvent and has been developed and licensed by Lurgi Oel Gas Chemie and Linde. The process is often applied for syngas purification and operates at temperatures as low as minus 30 to minus 100°F. Applications include selective or bulk removal of sour components. However the extremely low total sulfur impurities achievable with this process (down to less than 0.1 ppmv), are normally not required in natural gas service. Here the process is best suited where there are very high amounts of CO2 to be removed.
The process has been applied for the purification of natural gas for LNG production. However as pure cold methanol exhibits a certain co-absorption of higher hydrocarbons, implementation of the process in natural gas service has been limited to applications where only low concentrations of ethane and heavier components are present.
6.4.4 Ifpexol Process®
This process developed by IFP licensed by Prosernat uses refrigerated methanol solutions containing water in order to reduce hydrocarbon co-absorption. It is generally associated downstream of the Ifpex-1 dehydration process (Dehydration chapter) which simultaneously extracts heavier hydrocarbons. The Ifpex-2 process attains equivalent acid gas and mercaptans removal performance of the Rectisol process but at a slightly higher solvent rate due to the lower solvent purity.
6.4.5 Purisol®
This process has been developed and licensed by Lurgi Oel Gas Chemie. The solvent used is N-methyl-2-pyrrolidone (NMP or N-Pyrol), a high boiling point liquid. Purisol exhibits a selectivity for H2S, like Selexol, and features equivalent process possibilities.
As Purisol is also well suited for the selective removal of mercaptans; it can be used for the recovery of mercaptans from regeneration off gases from adsorptive mercaptan removal units.
6.4.5 Catasol®
This process is licensed by Eickmeyer and Associates. Catasol solvents are reported to have selectivity for H2S/CO2, CO2/propane, and COS/CO2.
6.4.6 Morphysorb®
A physical solvent process developed by Krupp Uhde called Morphysorb, uses N-formyl morpholine (NFM) as a physical solvent. The process claims lower investment and operating costs when processing gases heavily loaded with acid components.
6.5 Hybrid Solvent Processes
Table 6-4 showed the strengths and weaknesses of amine and physical solvent systems. To take advantage of the strengths of each type, a number of hybrid processes commercially used, and under development, combine physical solvents with amines. Depending upon the solvent−amine combination, nearly complete removal of H2S, CO2, and COS is possible. Other hybrid systems provide high H2S and COS removal while slipping CO2. Sulfinol® currently is one of the more commonly used processes. The process uses a combination of a physical solvent (sulfolane) with DIPA or MDEA. The selected amine depends upon the acid gases in the feed and whether CO2 removal is required. Like the physical solvent processes, the hybrid systems may absorb more hydrocarbons, including BTEX, but that property can be adjusted by varying water content.
6.5.1 Sulfinol® Process
The Sulfinol® Process, licensed by Shell Global Solutions, is used to remove H2S, CO2, COS, CS2, mercaptans and polysulfides from natural and synthetic gases. Sulfinol is a mixture of Sulfolane (a physical solvent), water and either DIPA or MDEA (both chemical solvents).
Solution concentrations range between 25% and 40% sulfolane, 40-55% DIPA, and 20-30% water and depend on the conditions and composition of the gas being treated. It is this dual capacity as both a physical and a chemical solvent that gives Sulfinol its advantages.
Sulfinol with DIPA (Sulfinol-D) is used when complete removal of H2S, CO2, and COS is desired. Sulfinol with MDEA (Sulfinol-M) is used for the selective removal of H2S in the presence of CO2, with partial removal of COS. Both Sulfinols can reduce the total sulfur content of treated gas down to low ppm levels. Some disadvantages are: a higher heavy hydrocarbon co-absorption, and a reclaimer is sometimes required for Sulfinol D when removing CO2.
The presence of the physical solvent, Sulfolane@, allows higher acid gas loadings compared to systems based on amine only. Typical loadings are 1.5 mol of acid gas per mole of Sulfinol solution. Higher acid gas loadings, together with a lower energy of regeneration, can result in lower capital and energy costs per unit of acid gas removed as compared to the ethanolamine processes.
Features of the Sulfinol process include essentially complete removal of mercaptans, high removal rate of COS, lower foaming tendency, lower corrosion rates, and the ability to slip up to 50% CO2.
The design is similar to that of the ethanolamines. Degradation of DIPA to oxazolidones (DIPA-OX) usually necessitates the installation of a reclaimer for their removal.
As with the ethanolamine processes, aromatics and heavy hydrocarbons in the feed gas should be removed prior to contact with the Sulfinol solution to minimize foaming.
The merits of the Sulfinol process, as compared to the ethanolamine processes, are many, but there are other factors that must be considered before selecting the appropriate gas treating process. For example, a licensing fee, while not necessary for the ethanolamine processes, is required for the Sulfinol process. The solvent costs are generally higher for the Sulfinol process than they are for DEA. Operators are more familiar with DEA and the typical problems associated with this process. In cases of low acid gas partial pressure, the advantage of a lower circulation rate for the Sulfinol process diminishes compared to the DEA process.
6.5.2 Selefining Process
This process developed by Snamprogetti uses tertiary amines such as dimethyl ethanol amine (DMEA) dissolved in an organic solvent with limited amounts of water. The process is very selective to H2S as CO2 hydration is almost completely avoided. It also removes other sulfur species such as mercaptans, COS and CS2. It has a tendency to co-absorb hydrocarbons which can to some extent be controlled by increasing the water content of the solvent.
Fundamentals of Oil and Gas Processing
Basics of Gas Field Processing
Basics of Corrosion in Oil and Gas Industry
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Basics of Gas Field Processing
Basics of Corrosion in Oil and Gas Industry
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Re: Basics of Gas Field Processing Book "Full text"
Gas Sweetening and Sulfur Recovery - Chapter 6 - Part 2 A
6.6 General Considerations for Solvent Process
6.6.1 Solution Filtration
Filtration of the treating solution to remove entrained solids is essential to the successful operation of a gas treating plant.
Filtration rates should be as high as practical and may range from 5 per cent of circulation to full stream. Removing particles down to 5 microns in size is recommended. In order to do this efficiently, two stages of filtration may be required. The first stage, typically a cartridge-type or precoat filter, is designed to remove particles down to the 10 micron or less range.
The second stage of filtration, typically an activated carbon filter, removes hydrocarbons and other contaminants. This is accomplished by adsorption. The carbon filter can also remove smaller particles from the amine stream. The carbon granule size can be selected to remove particles down to the 5 micron range. The activated carbon filter should always be located downstream of the first stage filter because the deposition of solids would plug the carbon filter.
The carryover of carbon fines can be controlled by either locating a second cartridge-type filter immediately downstream of the carbon filter or using a graded carbon bed. In a graded bed, larger granules are placed at the outlet of the filter to trap fines. Large carbon granules produce fewer fines but are less efficient for adsorption.
Basic degradation products are identified by gas chromatography and mass spectrometry. Acidic degradation products are identified by ion chromatography exclusion. These tests are recommended when the amine solution appears to lose its ability to pick up acid gas. Degradation products affect the results of the conventional estimation of amine concentration by titration. This may cause artificially high or low apparent amine concentrations Also, the carbon bed will adsorb very little strong acid degradation products. In this case, purging or reclamation of the solution is recommended.
Carbon filters can be partially regenerated with steam, which removes hydrocarbons and other adsorbed contaminants. Regeneration or bed change out is recommended when foam tests on the inlet and outlet streams show no improvement. This indicates carbon bed saturation.
Filters may be located on the lean or rich solution side. Filtration on the rich side seeks to remove particles that are more insoluble under the rich solution conditions. It also prevents solids accumulation in the hot environment of the stripper.
However, proper design and operating procedures for personnel protection during filter maintenance is mandatory when H2S may be present.
Filtration equipment should be used continuously beginning with the first day of plant operations. When starting up the plant, the full flow filter, even if temporary, may prove its worth by removing the scale and other solid particles and allowing much quicker and easier start-up of the plant.
6.6.2 Flash Tank
Rich solution leaving the contactor may pass through a flash tank. A flash tank is more important when treating high pressure gas. Gases entrained in the rich solution will be separated.
In addition, the amount of absorbed gas will be decreased because of the lower operating pressure of the flash tank. Using a flash tank will:
Reduce erosion in rich/lean exchangers.
Minimize the hydrocarbon content in the acid gas.
Reduce the vapor load on the stripper.
Possibly allow the off-gas from the flash tank to be used as fuel (may require sweetening).
When heavy hydrocarbons are present in the natural gas, the flash tank can also be used to skim off the heavy hydrocarbons that were absorbed by the solution. Residence times for flash tanks in amine service vary from 3-10 minutes depending on separation requirements. Inlet gas streams containing only methane and ethane require shorter residence times. Rich gas streams require longer time for the dissociation of gas from solution and/or the separation of liquid phases.
6.6.3 Corrosion
Corrosion is an operating concern in nearly all sweetening installations. The combination of H2S and CO2 with water practically ensures that corrosive conditions will exist in portions of the plant. In general, gas streams with high H2S to CO2 ratios are less corrosive than those having low H2S to CO2 ratios. H2S concentrations in the ppmv range with CO2 concentrations of 2 percent or more tend to be particularly corrosive.
Because the corrosion in sweetening plants tends to be chemical in nature, it is strongly a function of temperature and liquid velocity. The type of sweetening solution being used and the concentration of that solution has a strong impact on the corrosion rate. Increased corrosion can be expected with stronger solutions and higher gas loadings.
Hydrogen sulfide dissociates in water to form a weak acid. The acid attacks iron and forms insoluble iron sulfide. The iron sulfide will adhere to the base metal and may provide some protection from further corrosion, but it can be eroded away easily, exposing fresh metal for further attack. High liquid velocities can erode the protective iron sulfide film with resulting high corrosion rates.
CO2 in the presence of free water will form carbonic acid. The carbonic acid will attack iron to form a soluble iron bicarbonate which, upon heating, will release CO2 and an insoluble iron carbonate or hydrolize to iron oxide.
In general, design velocities in rich solution piping should be 50% of those that would be used in sweet service. Because of the temperature relationship to corrosion, the reboiler, the rich side of the amine-amine exchanger, tend to experience high corrosion rates. Because of the low pH the stripper overhead condensing loop also tends to experience high corrosion rates.
Acid degradation products also contribute to corrosion. A suggested mechanism for corrosion is that degradation products act as chelating agents for iron when hot. When cooled, the iron chelates become unstable, releasing the iron to form iron sulfide in the presence of H2S. Primary amines are thought to be more corrosive than secondary amines because the degradation products of the primary amines act as stronger chelating agents.
Several forms of stress corrosion cracking are possible in amine sweetening systems. Amine stress corrosion cracking can occur and is worst in hot solutions, but cracking can occur in cooler lines and both rich and lean streams. Wet sulfide cracking and blistering can occur due to hydrogen generated in corrosion reactions. The hydrogen can collect at small inclusions in the steel which delaminate and then link in a step-wise pattern to create blisters. This is called HIC or hydrogen induced cracking. Sometimes stress influences the cracking to cause SOHIC or stress oriented hydrogen induced cracking. HIC resistant steels are available. Seamless pipe is less prone to HIC than plate steels.
Corrosion in alkaline salt processes, such as the hot carbonate process, has been reported to range from none to severe. Corrosion can be expected where CO2 and steam are released through flashing. Severe erosion can take place when carbonate solution strengths exceed 40% because of the tendency to form bicarbonate crystals when the solution cools.
Many corrosion problems may be solved using corrosion inhibitors in combination with operating practices which reduce corrosion.
Following are some guidelines to minimize corrosion.
Maintain the lowest possible reboiler temperature.
If available, use low temperature heat medium rather than a high temperature heat medium or direct firing. When a high temperature heat medium or direct firing for the reboiler is used, caution should be taken to add only enough heat for stripping the solution.
Minimize solids and degradation products in the system through reclaimer operation and effective filtration.
Keep oxygen out of the system by providing a gas blanket on all storage tanks and maintain a positive pressure on the suction of all pumps.
Ensure deionized water or oxygen/chemical-free boiler condensate is used for make up water. If available, steam can be used to replace water loss.
Limit solution strengths to minimum levels required for treating.
Pipe solution exchangers for upflow operation with the rich solution on the tube side.
Monitor corrosion rates with coupons or suitable corrosion probes.
Maintain adequate solution level above reboiler tube bundles and fire tubes; a minimum tube submergence of 12" is recommended.
Corrosion inhibitors used include high molecular weight amines and heavy metal salts. These inhibitors offer potential savings in both capital and operating costs for these special cases.
6.6.4 Foaming
A sudden increase in differential pressure across a contactor or a sudden liquid level variation at the bottom of the contactor often indicates severe foaming. When foaming occurs, there is poor contact between the gas and the chemical solution. The result is reduced treating capacity and sweetening efficiency, possibly to the point that outlet specification cannot be met.
Some reasons for foaming are:
Suspended solids
Organic acids
Corrosion inhibitors
Condensed hydrocarbons
Soap-based valve greases
Makeup water impurities
Degradation products
Lube Oil
Foaming problems can usually be traced to plant operational problems. Contaminants from upstream operations can be minimized through adequate inlet separation. Condensation of hydrocarbons in the contactor can usually be avoided by maintaining the lean solution temperature at least 10°F above the hydrocarbon dew point temperature of the outlet gas.
Temporary upsets can be controlled by the addition of antifoam chemicals. These antifoams are usually of the silicone or long-chain alcohol type.
The following test for foaming should be run with the various types of inhibitors being considered for a given application. This test should give the operator an indication of which antifoam will be the most effective for the particular case. Place several drops of antifoam in 200 ml of treating solution contained in a 1000 ml cylinder. Bubble oil-free air through the solution at a constant rate. After five minutes have elapsed shut off the air and start a timer. Note the height of foam at the time the air was shut off and the amount of time required for the foam to break. The foam height is the difference between the height of the foam and the initial height of the liquid. The time for the foam to break is an indication of the stability of the foam. A comparison of antifoams will let the operator select which inhibitor will best solve his foaming problems. Between antifoam tests, care should be taken to clean the test cylinder thoroughly because a very small amount of inhibitor may affect the test.
6.6.5 Materials
Treating plants normally use carbon steel as the principal material of construction. Vessels and piping should be stress relieved in order to minimize stress corrosion along weld seams. Corrosion allowance for equipment ranges from 1/16" to 1/4", typically 1/8". In some instances, when corrosion is known to be a problem, or high solution loadings are required, stainless steel or clad stainless steel may be used in the following critical areas:
1. Reflux condenser
2. Reboiler tube bundle
3. Rich/lean exchanger tubes
4. Bubbling area of the contactor and/or stripper trays.
5. Rich solution piping from the rich/lean exchanger to the stripper.
6. Bottom 5 trays of the contactor and top 5 trays of stripper, if not all.
Usually 304, 316, or 410 stainless steel will be used in these areas, even through corrosion has been experienced with 410 stainless in DEA service for CO2 removal in the absence of H2S.
L grades are recommended if the alloys are to be welded.
Controlling oxygen content to less than 0.2 ppmw is effective in preventing chloride SCC in waters with up to 1000 ppmw chloride content, at temperatures up to 570°F. There has been an increased use of duplex stainless steels, and they have been successfully used in the water treatment industry to prevent chloride SCC in high chloride waters. This suggests duplex stainless steels could be utilized in amine plant service where high chloride content is expected. As with any specialty steel, proper fabrication techniques and welding procedures are required.
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6.7 Solid Bed Processes
6.7.1 General Process Description
A fixed bed of solid particles can be used to remove acid gases either through chemical reactions or through ionic bonding.
This process flows the gas stream through a fixed bed of solid particles, which removes the acid gases and holds them in the bed. When the bed is spent, the vessel must be removed from service and the bed regenerated or replaced. Since the bed must be removed from service to be regenerated, some spare capacity is normally provided.
Commonly used processes under this category are as follows:
Acid gas removal through chemical reactions include:
Iron Sponge Process, SulfaTreat®, Zinc Oxide Process, Chemsweet®, & PuraSpec®
Acid gas removal through Ionic bonding “Adsorption” include:
Molecular Sieve Process, & impregnated activated carbon.
6.7.2 Iron Sponge Process
The iron sponge process is economically applied to gases containing small amounts of H2S (<300 ppm) operating at low to moderate pressures in the range of 25-500 psig. This process does not remove CO2.
The reaction of iron oxide and H2S produces iron sulfide and water as follows:
Fe2O3 +3H2S → Fe2S3 +3H2O Eq. 6-12
FeO+H2S → FeS+H2O Eq. 6-13
The reaction requires the presence of slightly alkaline water (pH 8-10) and a temperature below 110 0F (47 0C). If the gas does not contain sufficient water vapor, water may need to be injected into the inlet gas stream. The pH level can be maintained through the injection of caustic soda, soda ash, lime, or ammonia with the water.
Although the presence of free alkalines enhances H2S removal, it also creates potential safety hazards and promotes the formation of undesirable salts, adding to capital costs.
Ferric oxide is impregnated on wood chips,which produce a solid bed with a large ferric oxide surface area. Several grades of treated wood chips are available, based on iron oxide content. Ferric oxide wood chips are available in 6.5, 9.0,15.0, and 20 lbs iron oxide/bushel.Chips are contained in a vessel, and sour gas flows downward through the bed and reacts with the ferric oxide. Figure 6-11 shows a vertical vessel used in the iron sponge process.
The bed can be regenerated with air; however, only about 60% of the previous bed life can be expected. The bed life of the batch process is dependent upon the quantity of H2S, the amount of iron oxide in the bed, residence time, pH, moisture content, and temperature.
Ferric sulfide can be oxidized with air to produce sulfur and regenerate the ferric oxide. Regeneration must be performed with care because the reaction with oxygen is exothermic (i.e., gives off heat). Air must be introduced slowly so that the heat of reaction can be dissipated.
If air is introduced quickly, the heat of reaction may ignite the bed.
For this reason, spent wood chips should be kept moist when removed from the vessel. Otherwise, the reaction with oxygen in the air may ignite the chips and cause them to smolder.
The reactions for oxygen regeneration are as follows:
2Fe2S3 + 3O2 + 2H2O → 2Fe2O3.(H2O) + 6S+Heat Eq. 6-14
4FeS +3O2 + 2xH2O → 2Fe2O3(H2O)x +4S +Heat Eq. 6-15
S2 +2O2 → 2SO2 Eq. 6-16
Some of the elemental sulfur produced in the regeneration step remains in the bed. After several cycles, this sulfur will cake over the ferric oxide, decreasing the reactivity of the bed and causing excessive gas pressure drop.
Typically after 10 cycles, the bed must be removed from the vessel and replaced with a new bed.
It is possible to operate an iron sponge with continuous regeneration by the introduction of small amounts of air in the sour gas feed. The oxygen in the air regenerates the iron sulfide and produces elemental sulfur. Although continuous regeneration decreases the amount of operating labor, it is not as effective as batch regeneration, and it may create an explosive mixture of air and natural gas. Due to the added costs associated with an air compressor, continuous regeneration generally does not prove to be the economic choice for the typically small quantities of gas involved.
Cooler operating temperatures of the natural gas, for example, during the winter, create the potential for hydrate formation in the iron sponge bed.
Hydrates can cause high-pressure drop, bed compaction, and flow channeling.
When the potential for hydrates exists, methanol can be injected to inhibit their formation. If insufficient water is present to absorb the methanol, it may coat the bed, forming undesirable salts. Hydrocarbon liquids in the gas tend to accumulate on the iron sponge media, thus inhibiting the reactions.
The use of a gas scrubber upstream of the iron sponge at a gas temperature slightly less than that of the sponge media may prevent significant quantities of liquids from condensing and fouling the bed.
There has been a recent revival in the use of iron sponges to sweeten light hydrocarbon liquids. Sour liquids flow through the bed and are contacted with the iron sponge media and the reaction proceeds as above.
Fig. 6-11 Iron oxide acid-gas treating unit.
6.7.3 SulfaTreat®
The Sulfa-Treat process similar to iron sponge process. It is economically applied to gases containing small amounts of H2S. This process utilizes a proprietary iron oxide co-product mixed with inert powder to form a porous bed. Sour gas flows through the bed and reacts with iron as in iron-sponge reaction.
The material of SulfaTreat is a dry, free-flowing granular substance used for selective removal of H2S and mercaptans from natural gas in the presence of CO2. It is not affected by CO2, and it does not produce sulfur or nitric oxide (NOx). Also SulfaTreat will not ignite or "cement up" in the vessel. Other advantages include longer bed life and lower cost.
SulfaTreat’s particle size varies from 4 to 30 mesh and has a bulk density of 70 lb/ft3. These physical properties give uniform porosity and permeability, which offers small resistance to flow and resistance to bed compression at normal velocities.
Applications for SulfaTreat include: natural gas treating, amine treater off gas, high concentration CO2 streams, and any other H2S-containing system.
The reaction works better with saturated gas and at elevated temperature up to 1300F (54.40C). There is no minimum moisture or pH level required. The amount of bed volume required increases as the velocity increases and as the bed height decreases. Operation of the system below 400F (4.40C) is not recommended. Beds are not regenerated and must be replaced when the bed is spent.
6.7.4 Zinc Oxide Process
The equipment used in the zinc oxide process is similar to the iron sponge process. It uses a solid bed of granular zinc oxide to react with the H2S to form zinc sulfide and water as shown below
ZnO+H2S → ZnS+H2O Eq. 6-17
The rate of reaction is controlled by the diffusion process, as the sulfide ion must first diffuse to the surface of the zinc oxide to react. Temperatures above 2500F (1200C) increase the diffusion rate, which promotes the reaction rate. The strong dependence on diffusion means that other variables, such as pressure and gas velocity, have little effect on the reaction.
The zinc oxide is contained in long thin beds to lessen the chances of channeling. The pressure drop through the beds is low. Bed life is a function of gas H2S content and can vary from 6 months to more than 10 years. Beds are often used in series to increase the level of saturation prior to change out of the catalyst. A spent bed is discharged by gravity flow through the bottom of the vessel.
The zinc oxide process is seldom used due to disposal problems with the spent bed, which is classified as a heavy metal salt.
6.7.5 Chemsweet®
Chemsweet® is the name for another batch process for the removal of H2S from natural gas. Chemicals used are a mixture of zinc oxide, zinc acetate, water, and a dispersant to keep solid particles in suspension. Natural gas is bubbled through the solution where H2S reacts with zinc oxide. Though several reactions take place in solution, the net result is that zinc oxide reacts with H2S to form zinc sulfide and water.
Chemsweet® can treat gas streams with H2S concentration up to 400 ppmv. and has been operated between pressures of 75 and 1,400 psig. Mercaptan concentrations in excess of 10% of the H2S concentration in the gas stream can cause a problem. Some of the mercaptans will react with the zinc oxide and be removed from the gas. The resulting zinc mercaptides [Zn(OH)RH] will form a sludge and possibly cause foaming problems.
6.7.6 PuraSpec®
Johnson Matthey Catalysts supplies the PuraSpec range of processes and products for desulfurization of hydrocarbon gases and liquids. The processes use fixed beds of granular, metal oxide-based chemical absorbents which are developments of the ‘high temperature zinc oxide’ used for purification of hydrocarbon feedstocks to steam reformers in ammonia, hydrogen, and methanol plants. PuraSpec absorbents are effective at temperatures down to 32°F, so no added heat is necessary, and are in service at pressures from atmospheric, treating vent gases, to 1800 psi treating dense phase gas feed to a gas processing plant.
PuraSpec units are in service treating natural gas to pipeline or petrochemical specifications. Because the absorbents remove H2S and COS irreversibly, they are best suited to polishing duties.
6.7.8 Molecular Sieve Process
Acid gases, as well as water, can be effectively removed by physical adsorption on synthetic zeolites. Applications of acid gas removal using molecular sieve are limited because water displaces acid gases on the adsorbent bed. (Chapter 5 provides more details on adsorption and its use in dehydration.)
Crystalline structure of the solids provides a very porous material having uniform pore size. Within the pores the crystalline structure creates a large number of localized polar charges called active sites. Polar gas molecules, such as H2S and water vapor, which enter the pores, form weak ionic bonds at the active sites. Nonpolar molecules, such as paraffin hydrocarbons, will not bond to the active sites.
Carbon dioxide molecules are about the same size as H2S molecules, but are nonpolar. CO2 will enter the pores but will not bond to the active sites. Small quantities of CO2 will be removed by becoming trapped in the pores by bonded H2S or H2O molecules blocking the pores.
CO2 will obstruct the access of H2S and H2O to the active sites, thus decreasing the overall effectiveness of the molecular sieve. Beds must be sized to remove all H2O and provide for interference from other molecules in order to remove all H2S.
The adsorption process usually occurs at moderate pressure. Ionic bonds tend to achieve an optimum performance near 450 psig but can operate in a wide range of pressures.
Hydrogen sulfide can be selectively removed to meet 0.25 grain/100 scf (4ppm) specification. However, this reduction requires regeneration of the bed at 600°F (315°C) for extended time (Usually one hour or more depending upon process conditions) with the potential for COS formation if 4A is used (Molecular sieves are classified by their nominal pore diameter in Angstroms. See Chapter 5 for details.).
The problem of COS formation during processing according to the reaction in Equation 6-18, has been extensively studied.
H2S + CO2 ↔ COS + H2O Eq. 6-18
Molecular sieve products have been developed that do not catalyze COS formation. The central zone in the regeneration cycle is most favorable to COS formation.
Regeneration of a molecular sieve bed concentrates the H2S into a small regeneration stream which must be treated or disposed of.
Figure 6-12 shows a typical flow diagram for removal of H2S from natural gas. The configuration is similar to that for dehydration but with the significant difference that the regeneration gas contains high quantities of H2S as well as water as it leaves the adsorbent bed and, thus, must be treated. (Chapter 5 presents a thorough discussion of adsorption.) The flow configuration shows the first bed in the adsorption cycle, the second bed cooling down after regeneration, and the third bed undergoing regeneration with hot gas.
Fig. 6-12. Integrated Natural Gas Desulfurization Plant
Chi and Lee (1973) studied the coadsorption of H2S, CO2, and H2O on a 5A molecular sieve from a natural gas mixture under a variety of conditions.
Figure 6-13, from their paper, shows a typical concentration versus time curve.
In the figure, y is the concentration in the exit stream and yo is the concentration in the inlet to the bed. The gas that entered the bed was saturated with H2O and contained both CO2 (1.14 mol%) and H2S (0.073 mol%.). Because the CO2 content of the gas was 15.6 times that of the H2S, the bed quickly saturated with CO2, and its breakthrough was almost instantaneous. As the H2S was adsorbed and moved down the column, it displaced the CO2 and, consequently, after approximately 30 minutes, the CO2 exit concentration peaked at a value greater than its inlet concentration.
The same phenomenon occurs when the H2S is displaced by the water.
Because the H2S must be removed to extremely low levels, the bed is effectively exhausted from the perspective of H2S purification shortly after H2S breakthrough occurs and, thus, well before the bed is totally saturated.
Fig. 6.13 Effluent H2S and CO2 concentration from adsorption bed as a function of Time.
In general, the sieve bed can be designed to dehydrate and sweeten simultaneously.
The Engineering Data Book (GPSA) notes that a key point in adsorber design is to properly design for treatment of the regeneration gas because the peak H2S concentration may be 30 times the H2S concentration in the feed.
Care should be taken to minimize mechanical damage to the solid crystals as this will decrease the bed’s effectiveness. The main cause of mechanical degradation is the sudden pressure and/or temperature changes that may occur when switching from adsorption to regeneration cycles. Proper instrumentation can significantly extend bed life.
The molecular sieve process is limited to small gas streams operating at moderate pressures. It is generally used for polishing applications following one of the other processes.
6.7.9 Oxorbon
Oxorbon is an alternative solid bed material which consists of activated carbon impregnated with potassium iodide (KI). Donau Carbon markets such a carbon for the removal of H2S and mercaptans. The adsorbed H2S is converted to elemental sulfur by catalytic reaction under the presence of oxygen. The resulting sulfur is fixed on the pores of the activated carbon. Sulfur loadings as high as 60% of the carbon mass have been reported and sulfide concentrations below 1 ppmv are claimed.
6.8 Direct Conversion Processes (Liquid Redox)
Chemical and physical solvent processes remove acid gas from the natural gas stream but release H2S and CO2 when the solvent is regenerated. The release of H2S to the atmosphere is limited by environmental regulations. Acid gases could be routed to an incinerator/flare, which would convert the H2S to SO2. Environmental regulations restrict the amount of SO2 vented or flared. Direct conversion processes use chemical reactions to oxidize H2S and produce elemental sulfur. These processes are generally based either on the reaction of H2S and O2 or H2S and SO2. Both reactions yield water and elemental sulfur. These processes are licensed and involve specialized catalysts and/or solvents.
The redox agent is then regenerated by reaction with air in an oxidizer vessel.
Liquid oxidation−reduction (redox) processes use iron as an oxidizing agent in solution. The process involves four basic steps:
1. Removal of H2S from the gas by absorption into a caustic solution
2. Oxidation of the HS− ion to elemental sulfur via the oxidizing agent
3. Separation and removal of sulfur from the solution
4. Regeneration (i.e., oxidation) of the oxidizing agent by use of air
The processes share the general chemistry shown below.
• H2S absorption using an alkaline solution:
H2S (g) ↔ H2S (sol’n) Eq 6-19
H2S (sol’n) ↔ H+ + HS− Eq 6-20
Note that this is a non-selective, basic scrubbing solution; thus, additional acidic components (CO2 , HCN) of the sour gas will be soluble in the scrubbing solution. While CO2 is partially absorbed in the alkaline solution, it does not participate in the redox reactions. CO2 may however increase the rate of consumption of caustic or buffering compounds. The fate of other reduced sulfur species (COS, CS2, RSH, RSR) depends on the particular process being considered.
• Conversion to Elemental Sulfur:
HS− + H+ + 1/2 O2 (sol’n) → H2O (l) + S Eq 6-21
Note that in the absence of an auxiliary redox reagent (ARR) the reaction shown as Eq 21-24 is slow and nonspecific. The addition of an (ARR) increases the rate of reaction and directs the oxidation to elemental sulfur.
HS− + ARR(OX) → S + H+ + ARR (RED) Eq 6-22
where (OX) denotes the oxidized form and (RED) denotes the reduced form of the ARR.
• Regeneration of the Spent ARR Using Air:
O2 (g) ↔ O2 (sol’n) Eq 6-23
ARR (RED) + 1/2 O2 (sol’n) + 2 H+ → ARR(OX) + H2O(l) Eq 6-24
The overall simplified chemistry of liquid redox processes is thus:
H2S(g) + 1/2 O2 (g) → H2O(l) + S Eq 6-25
While the oxidizing agent forms the desired reaction, it also reacts with the sulfur species to form metal sulfides that precipitate from solution. To avoid the metal sulfide reaction, the solution contains chelating agents (organic compounds that bind with the metal ion to restrict its reactivity but still permit electron transfer for oxidation−reduction reactions). A common chelating agent for iron is EDTA (ethylenediaminetetraacetic acid). Several processes, including Lo-Cat II®, SulFerox ®, and Sulfint-HP, use iron and chelating agents.
6.8.1 Stretford Process
The Stretford process involves the use of vanadium salts as the ARR. The process has been extensively used in Europe. However environmental concerns around the discharge of vanadium compounds has limited its use.
Figure 6.14 shows a simplified diagram of the Stretford process. The gas stream is washed with an aqueous solution of sodium carbonate, sodium vanadate, and anthraquinone disulfonic. An oxidized solution is delivered from the pumping tank to the top of the absorber tower where it contacts the gas stream in a counter-current flow.
The bottom of the absorber tower consists of a reaction tank from which the reduced solution passes to the solution flash drum, which is situated above the oxidizer. The reduced solution passes from here into the base of the oxidizer vessel. Hydrocarbon gases, which have been dissolved in the solution at the plant pressure, are released from the top of the flash drum.
Air is blown into the oxidizer, and the main body of the solution, now reoxidized, passes into the pumping tank. The sulfur is carried to the top of the oxidizer by froth created by the aeration of the solution and passes into the thickener.
The function of the thickener is to increase the weight percent of sulfur that is pumped to one of the alternate sulfur recovery methods of filtration, filtration and autoclaves, centrifugation or centrifugation with heating.
Chemical reactions involved are:
H2S+Na2CO3 → NaHS+NaHCO3 Eq 6-26
Sodium carbonate provides the alkaline solution for initial adsorption of H2S and the formation of hydrosulfide (HS). The hydrosulfide is reduced in a reaction with sodium meta vanadate to precipitate sulfur
HS- +V+5 → S +V+4 Eq 6-27
Anthraquinone disulfonic acid (ADA) reacts with 4-valent vanadium and converts it back to 5 valent
V+4 +ADA → V+5 +ADA (reduced) Eq 6-28
Oxygen from the air converts the reduced ADA back to the oxidized state as shown below:
Reduced ADA+O2 → ADA+H2O Eq 6-29
The overall reaction is
2H2S+O2 → 2H2O+2S Eq 6-30
6.8.2 Lo-Cat process
This is a liquid-redox process. The LO-CAT® process is a patented, wet scrubbing, liquid redox system that uses a chelated iron solution to convert H2S to innocuous, elemental sulfur. It does not use any toxic chemicals and does not produce any hazardous waste byproducts. In general, the LO-CAT process can directly treat a gas stream, or treat the H2S containing stream from an Acid Gas Removal unit, although the direct treat capability is limited to low pressure streams.
Processes employ high iron concentration reduction-oxidation technology for the selective removal of H2S (not reactive to CO2) to <4 ppm in both low and high pressure gas streams. Figure 6-15 shows a simplified flow schematic of the LO-CAT process.
Acid gas stream is contacted with the solution where H2S reacts with, and reduces the chelated iron and produces elemental sulfur. The iron is then regenerated by bubbling air through the solution. Heat is not required for regeneration. The reactions involved are exothermic (give off heat):
Absorption/reduction: 2Fe3+ +H2S → 2Fe2+ +S+2H+ Eq 6-31
Regeneration/oxidation: 2Fe2+ +½ O2 +2H+ → 2Fe3+ +H2O Eq 6-32
Overall chemistry: H2S + ½ O2 → S + H2O Eq 6-33
Fig 6.14. Simplified flow schematic of the Stretford process.
Fig. 6.15. Simplified flow schematic of the LO-CAT process.
6.8.3 Sulferox process
SulFerox is a Shell proprietary Iron Redox process, whereby a sour gas stream, containing hydrogen sulphide, is contacted with a liquid, containing soluble ferric (Fe3+) ions. In the process the H2S is oxidized to elemental sulphur and the Fe3+ is reduced to ferrous (Fe2+) ions. The system is regenerable, Fe2+ is subsequently reconverted to Fe3+ by oxidation with air. Sulphur is recovered from the aqueous solution as a moist cake.
The process offers low capital and operating costs through the use of a high concentration iron chelate solution and effective control of chelate degradation. The process also offers a patented contactor design to improve the overall efficiency of the process which further reduces capital costs. The optimum application for SulFerox is in the one to twenty tons per day range of recovered sulfur.
Typical Sulferox process equipment is illustrated in Figure 6.16. In this process, the sour gas is contacted in a small contactor in co-current flow with a water solution containing about 4% ferric iron ions, held in solution by proprietary chelate ligand. The H2S ionizes in the solution, and the ferric ions exchange electrons with the sulfur ions to form ferrous ions and elemental sulfur. The gas and the solution leave the contactor and are flowed into a separator. The gas out of the separator is sweet and requires further treating for dewpoint control. The solution is flowed into additional vessels for separating the elemental sulfur and for restoring the ferrous iron ions back to the active ferric iron state by contacting the solution with air.
In general, the Sulferox process can directly treat a gas stream, or treat the H2S containing stream from an Acid Gas Removal unit, although the direct treat capability is limited to low pressure streams.
Fig. 6.16. Simplified flow schematic of the sulferox process.
6.6 General Considerations for Solvent Process
6.6.1 Solution Filtration
Filtration of the treating solution to remove entrained solids is essential to the successful operation of a gas treating plant.
Filtration rates should be as high as practical and may range from 5 per cent of circulation to full stream. Removing particles down to 5 microns in size is recommended. In order to do this efficiently, two stages of filtration may be required. The first stage, typically a cartridge-type or precoat filter, is designed to remove particles down to the 10 micron or less range.
The second stage of filtration, typically an activated carbon filter, removes hydrocarbons and other contaminants. This is accomplished by adsorption. The carbon filter can also remove smaller particles from the amine stream. The carbon granule size can be selected to remove particles down to the 5 micron range. The activated carbon filter should always be located downstream of the first stage filter because the deposition of solids would plug the carbon filter.
The carryover of carbon fines can be controlled by either locating a second cartridge-type filter immediately downstream of the carbon filter or using a graded carbon bed. In a graded bed, larger granules are placed at the outlet of the filter to trap fines. Large carbon granules produce fewer fines but are less efficient for adsorption.
Basic degradation products are identified by gas chromatography and mass spectrometry. Acidic degradation products are identified by ion chromatography exclusion. These tests are recommended when the amine solution appears to lose its ability to pick up acid gas. Degradation products affect the results of the conventional estimation of amine concentration by titration. This may cause artificially high or low apparent amine concentrations Also, the carbon bed will adsorb very little strong acid degradation products. In this case, purging or reclamation of the solution is recommended.
Carbon filters can be partially regenerated with steam, which removes hydrocarbons and other adsorbed contaminants. Regeneration or bed change out is recommended when foam tests on the inlet and outlet streams show no improvement. This indicates carbon bed saturation.
Filters may be located on the lean or rich solution side. Filtration on the rich side seeks to remove particles that are more insoluble under the rich solution conditions. It also prevents solids accumulation in the hot environment of the stripper.
However, proper design and operating procedures for personnel protection during filter maintenance is mandatory when H2S may be present.
Filtration equipment should be used continuously beginning with the first day of plant operations. When starting up the plant, the full flow filter, even if temporary, may prove its worth by removing the scale and other solid particles and allowing much quicker and easier start-up of the plant.
6.6.2 Flash Tank
Rich solution leaving the contactor may pass through a flash tank. A flash tank is more important when treating high pressure gas. Gases entrained in the rich solution will be separated.
In addition, the amount of absorbed gas will be decreased because of the lower operating pressure of the flash tank. Using a flash tank will:
Reduce erosion in rich/lean exchangers.
Minimize the hydrocarbon content in the acid gas.
Reduce the vapor load on the stripper.
Possibly allow the off-gas from the flash tank to be used as fuel (may require sweetening).
When heavy hydrocarbons are present in the natural gas, the flash tank can also be used to skim off the heavy hydrocarbons that were absorbed by the solution. Residence times for flash tanks in amine service vary from 3-10 minutes depending on separation requirements. Inlet gas streams containing only methane and ethane require shorter residence times. Rich gas streams require longer time for the dissociation of gas from solution and/or the separation of liquid phases.
6.6.3 Corrosion
Corrosion is an operating concern in nearly all sweetening installations. The combination of H2S and CO2 with water practically ensures that corrosive conditions will exist in portions of the plant. In general, gas streams with high H2S to CO2 ratios are less corrosive than those having low H2S to CO2 ratios. H2S concentrations in the ppmv range with CO2 concentrations of 2 percent or more tend to be particularly corrosive.
Because the corrosion in sweetening plants tends to be chemical in nature, it is strongly a function of temperature and liquid velocity. The type of sweetening solution being used and the concentration of that solution has a strong impact on the corrosion rate. Increased corrosion can be expected with stronger solutions and higher gas loadings.
Hydrogen sulfide dissociates in water to form a weak acid. The acid attacks iron and forms insoluble iron sulfide. The iron sulfide will adhere to the base metal and may provide some protection from further corrosion, but it can be eroded away easily, exposing fresh metal for further attack. High liquid velocities can erode the protective iron sulfide film with resulting high corrosion rates.
CO2 in the presence of free water will form carbonic acid. The carbonic acid will attack iron to form a soluble iron bicarbonate which, upon heating, will release CO2 and an insoluble iron carbonate or hydrolize to iron oxide.
In general, design velocities in rich solution piping should be 50% of those that would be used in sweet service. Because of the temperature relationship to corrosion, the reboiler, the rich side of the amine-amine exchanger, tend to experience high corrosion rates. Because of the low pH the stripper overhead condensing loop also tends to experience high corrosion rates.
Acid degradation products also contribute to corrosion. A suggested mechanism for corrosion is that degradation products act as chelating agents for iron when hot. When cooled, the iron chelates become unstable, releasing the iron to form iron sulfide in the presence of H2S. Primary amines are thought to be more corrosive than secondary amines because the degradation products of the primary amines act as stronger chelating agents.
Several forms of stress corrosion cracking are possible in amine sweetening systems. Amine stress corrosion cracking can occur and is worst in hot solutions, but cracking can occur in cooler lines and both rich and lean streams. Wet sulfide cracking and blistering can occur due to hydrogen generated in corrosion reactions. The hydrogen can collect at small inclusions in the steel which delaminate and then link in a step-wise pattern to create blisters. This is called HIC or hydrogen induced cracking. Sometimes stress influences the cracking to cause SOHIC or stress oriented hydrogen induced cracking. HIC resistant steels are available. Seamless pipe is less prone to HIC than plate steels.
Corrosion in alkaline salt processes, such as the hot carbonate process, has been reported to range from none to severe. Corrosion can be expected where CO2 and steam are released through flashing. Severe erosion can take place when carbonate solution strengths exceed 40% because of the tendency to form bicarbonate crystals when the solution cools.
Many corrosion problems may be solved using corrosion inhibitors in combination with operating practices which reduce corrosion.
Following are some guidelines to minimize corrosion.
Maintain the lowest possible reboiler temperature.
If available, use low temperature heat medium rather than a high temperature heat medium or direct firing. When a high temperature heat medium or direct firing for the reboiler is used, caution should be taken to add only enough heat for stripping the solution.
Minimize solids and degradation products in the system through reclaimer operation and effective filtration.
Keep oxygen out of the system by providing a gas blanket on all storage tanks and maintain a positive pressure on the suction of all pumps.
Ensure deionized water or oxygen/chemical-free boiler condensate is used for make up water. If available, steam can be used to replace water loss.
Limit solution strengths to minimum levels required for treating.
Pipe solution exchangers for upflow operation with the rich solution on the tube side.
Monitor corrosion rates with coupons or suitable corrosion probes.
Maintain adequate solution level above reboiler tube bundles and fire tubes; a minimum tube submergence of 12" is recommended.
Corrosion inhibitors used include high molecular weight amines and heavy metal salts. These inhibitors offer potential savings in both capital and operating costs for these special cases.
6.6.4 Foaming
A sudden increase in differential pressure across a contactor or a sudden liquid level variation at the bottom of the contactor often indicates severe foaming. When foaming occurs, there is poor contact between the gas and the chemical solution. The result is reduced treating capacity and sweetening efficiency, possibly to the point that outlet specification cannot be met.
Some reasons for foaming are:
Suspended solids
Organic acids
Corrosion inhibitors
Condensed hydrocarbons
Soap-based valve greases
Makeup water impurities
Degradation products
Lube Oil
Foaming problems can usually be traced to plant operational problems. Contaminants from upstream operations can be minimized through adequate inlet separation. Condensation of hydrocarbons in the contactor can usually be avoided by maintaining the lean solution temperature at least 10°F above the hydrocarbon dew point temperature of the outlet gas.
Temporary upsets can be controlled by the addition of antifoam chemicals. These antifoams are usually of the silicone or long-chain alcohol type.
The following test for foaming should be run with the various types of inhibitors being considered for a given application. This test should give the operator an indication of which antifoam will be the most effective for the particular case. Place several drops of antifoam in 200 ml of treating solution contained in a 1000 ml cylinder. Bubble oil-free air through the solution at a constant rate. After five minutes have elapsed shut off the air and start a timer. Note the height of foam at the time the air was shut off and the amount of time required for the foam to break. The foam height is the difference between the height of the foam and the initial height of the liquid. The time for the foam to break is an indication of the stability of the foam. A comparison of antifoams will let the operator select which inhibitor will best solve his foaming problems. Between antifoam tests, care should be taken to clean the test cylinder thoroughly because a very small amount of inhibitor may affect the test.
6.6.5 Materials
Treating plants normally use carbon steel as the principal material of construction. Vessels and piping should be stress relieved in order to minimize stress corrosion along weld seams. Corrosion allowance for equipment ranges from 1/16" to 1/4", typically 1/8". In some instances, when corrosion is known to be a problem, or high solution loadings are required, stainless steel or clad stainless steel may be used in the following critical areas:
1. Reflux condenser
2. Reboiler tube bundle
3. Rich/lean exchanger tubes
4. Bubbling area of the contactor and/or stripper trays.
5. Rich solution piping from the rich/lean exchanger to the stripper.
6. Bottom 5 trays of the contactor and top 5 trays of stripper, if not all.
Usually 304, 316, or 410 stainless steel will be used in these areas, even through corrosion has been experienced with 410 stainless in DEA service for CO2 removal in the absence of H2S.
L grades are recommended if the alloys are to be welded.
Controlling oxygen content to less than 0.2 ppmw is effective in preventing chloride SCC in waters with up to 1000 ppmw chloride content, at temperatures up to 570°F. There has been an increased use of duplex stainless steels, and they have been successfully used in the water treatment industry to prevent chloride SCC in high chloride waters. This suggests duplex stainless steels could be utilized in amine plant service where high chloride content is expected. As with any specialty steel, proper fabrication techniques and welding procedures are required.
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6.7 Solid Bed Processes
6.7.1 General Process Description
A fixed bed of solid particles can be used to remove acid gases either through chemical reactions or through ionic bonding.
This process flows the gas stream through a fixed bed of solid particles, which removes the acid gases and holds them in the bed. When the bed is spent, the vessel must be removed from service and the bed regenerated or replaced. Since the bed must be removed from service to be regenerated, some spare capacity is normally provided.
Commonly used processes under this category are as follows:
Acid gas removal through chemical reactions include:
Iron Sponge Process, SulfaTreat®, Zinc Oxide Process, Chemsweet®, & PuraSpec®
Acid gas removal through Ionic bonding “Adsorption” include:
Molecular Sieve Process, & impregnated activated carbon.
6.7.2 Iron Sponge Process
The iron sponge process is economically applied to gases containing small amounts of H2S (<300 ppm) operating at low to moderate pressures in the range of 25-500 psig. This process does not remove CO2.
The reaction of iron oxide and H2S produces iron sulfide and water as follows:
Fe2O3 +3H2S → Fe2S3 +3H2O Eq. 6-12
FeO+H2S → FeS+H2O Eq. 6-13
The reaction requires the presence of slightly alkaline water (pH 8-10) and a temperature below 110 0F (47 0C). If the gas does not contain sufficient water vapor, water may need to be injected into the inlet gas stream. The pH level can be maintained through the injection of caustic soda, soda ash, lime, or ammonia with the water.
Although the presence of free alkalines enhances H2S removal, it also creates potential safety hazards and promotes the formation of undesirable salts, adding to capital costs.
Ferric oxide is impregnated on wood chips,which produce a solid bed with a large ferric oxide surface area. Several grades of treated wood chips are available, based on iron oxide content. Ferric oxide wood chips are available in 6.5, 9.0,15.0, and 20 lbs iron oxide/bushel.Chips are contained in a vessel, and sour gas flows downward through the bed and reacts with the ferric oxide. Figure 6-11 shows a vertical vessel used in the iron sponge process.
The bed can be regenerated with air; however, only about 60% of the previous bed life can be expected. The bed life of the batch process is dependent upon the quantity of H2S, the amount of iron oxide in the bed, residence time, pH, moisture content, and temperature.
Ferric sulfide can be oxidized with air to produce sulfur and regenerate the ferric oxide. Regeneration must be performed with care because the reaction with oxygen is exothermic (i.e., gives off heat). Air must be introduced slowly so that the heat of reaction can be dissipated.
If air is introduced quickly, the heat of reaction may ignite the bed.
For this reason, spent wood chips should be kept moist when removed from the vessel. Otherwise, the reaction with oxygen in the air may ignite the chips and cause them to smolder.
The reactions for oxygen regeneration are as follows:
2Fe2S3 + 3O2 + 2H2O → 2Fe2O3.(H2O) + 6S+Heat Eq. 6-14
4FeS +3O2 + 2xH2O → 2Fe2O3(H2O)x +4S +Heat Eq. 6-15
S2 +2O2 → 2SO2 Eq. 6-16
Some of the elemental sulfur produced in the regeneration step remains in the bed. After several cycles, this sulfur will cake over the ferric oxide, decreasing the reactivity of the bed and causing excessive gas pressure drop.
Typically after 10 cycles, the bed must be removed from the vessel and replaced with a new bed.
It is possible to operate an iron sponge with continuous regeneration by the introduction of small amounts of air in the sour gas feed. The oxygen in the air regenerates the iron sulfide and produces elemental sulfur. Although continuous regeneration decreases the amount of operating labor, it is not as effective as batch regeneration, and it may create an explosive mixture of air and natural gas. Due to the added costs associated with an air compressor, continuous regeneration generally does not prove to be the economic choice for the typically small quantities of gas involved.
Cooler operating temperatures of the natural gas, for example, during the winter, create the potential for hydrate formation in the iron sponge bed.
Hydrates can cause high-pressure drop, bed compaction, and flow channeling.
When the potential for hydrates exists, methanol can be injected to inhibit their formation. If insufficient water is present to absorb the methanol, it may coat the bed, forming undesirable salts. Hydrocarbon liquids in the gas tend to accumulate on the iron sponge media, thus inhibiting the reactions.
The use of a gas scrubber upstream of the iron sponge at a gas temperature slightly less than that of the sponge media may prevent significant quantities of liquids from condensing and fouling the bed.
There has been a recent revival in the use of iron sponges to sweeten light hydrocarbon liquids. Sour liquids flow through the bed and are contacted with the iron sponge media and the reaction proceeds as above.
Fig. 6-11 Iron oxide acid-gas treating unit.
6.7.3 SulfaTreat®
The Sulfa-Treat process similar to iron sponge process. It is economically applied to gases containing small amounts of H2S. This process utilizes a proprietary iron oxide co-product mixed with inert powder to form a porous bed. Sour gas flows through the bed and reacts with iron as in iron-sponge reaction.
The material of SulfaTreat is a dry, free-flowing granular substance used for selective removal of H2S and mercaptans from natural gas in the presence of CO2. It is not affected by CO2, and it does not produce sulfur or nitric oxide (NOx). Also SulfaTreat will not ignite or "cement up" in the vessel. Other advantages include longer bed life and lower cost.
SulfaTreat’s particle size varies from 4 to 30 mesh and has a bulk density of 70 lb/ft3. These physical properties give uniform porosity and permeability, which offers small resistance to flow and resistance to bed compression at normal velocities.
Applications for SulfaTreat include: natural gas treating, amine treater off gas, high concentration CO2 streams, and any other H2S-containing system.
The reaction works better with saturated gas and at elevated temperature up to 1300F (54.40C). There is no minimum moisture or pH level required. The amount of bed volume required increases as the velocity increases and as the bed height decreases. Operation of the system below 400F (4.40C) is not recommended. Beds are not regenerated and must be replaced when the bed is spent.
6.7.4 Zinc Oxide Process
The equipment used in the zinc oxide process is similar to the iron sponge process. It uses a solid bed of granular zinc oxide to react with the H2S to form zinc sulfide and water as shown below
ZnO+H2S → ZnS+H2O Eq. 6-17
The rate of reaction is controlled by the diffusion process, as the sulfide ion must first diffuse to the surface of the zinc oxide to react. Temperatures above 2500F (1200C) increase the diffusion rate, which promotes the reaction rate. The strong dependence on diffusion means that other variables, such as pressure and gas velocity, have little effect on the reaction.
The zinc oxide is contained in long thin beds to lessen the chances of channeling. The pressure drop through the beds is low. Bed life is a function of gas H2S content and can vary from 6 months to more than 10 years. Beds are often used in series to increase the level of saturation prior to change out of the catalyst. A spent bed is discharged by gravity flow through the bottom of the vessel.
The zinc oxide process is seldom used due to disposal problems with the spent bed, which is classified as a heavy metal salt.
6.7.5 Chemsweet®
Chemsweet® is the name for another batch process for the removal of H2S from natural gas. Chemicals used are a mixture of zinc oxide, zinc acetate, water, and a dispersant to keep solid particles in suspension. Natural gas is bubbled through the solution where H2S reacts with zinc oxide. Though several reactions take place in solution, the net result is that zinc oxide reacts with H2S to form zinc sulfide and water.
Chemsweet® can treat gas streams with H2S concentration up to 400 ppmv. and has been operated between pressures of 75 and 1,400 psig. Mercaptan concentrations in excess of 10% of the H2S concentration in the gas stream can cause a problem. Some of the mercaptans will react with the zinc oxide and be removed from the gas. The resulting zinc mercaptides [Zn(OH)RH] will form a sludge and possibly cause foaming problems.
6.7.6 PuraSpec®
Johnson Matthey Catalysts supplies the PuraSpec range of processes and products for desulfurization of hydrocarbon gases and liquids. The processes use fixed beds of granular, metal oxide-based chemical absorbents which are developments of the ‘high temperature zinc oxide’ used for purification of hydrocarbon feedstocks to steam reformers in ammonia, hydrogen, and methanol plants. PuraSpec absorbents are effective at temperatures down to 32°F, so no added heat is necessary, and are in service at pressures from atmospheric, treating vent gases, to 1800 psi treating dense phase gas feed to a gas processing plant.
PuraSpec units are in service treating natural gas to pipeline or petrochemical specifications. Because the absorbents remove H2S and COS irreversibly, they are best suited to polishing duties.
6.7.8 Molecular Sieve Process
Acid gases, as well as water, can be effectively removed by physical adsorption on synthetic zeolites. Applications of acid gas removal using molecular sieve are limited because water displaces acid gases on the adsorbent bed. (Chapter 5 provides more details on adsorption and its use in dehydration.)
Crystalline structure of the solids provides a very porous material having uniform pore size. Within the pores the crystalline structure creates a large number of localized polar charges called active sites. Polar gas molecules, such as H2S and water vapor, which enter the pores, form weak ionic bonds at the active sites. Nonpolar molecules, such as paraffin hydrocarbons, will not bond to the active sites.
Carbon dioxide molecules are about the same size as H2S molecules, but are nonpolar. CO2 will enter the pores but will not bond to the active sites. Small quantities of CO2 will be removed by becoming trapped in the pores by bonded H2S or H2O molecules blocking the pores.
CO2 will obstruct the access of H2S and H2O to the active sites, thus decreasing the overall effectiveness of the molecular sieve. Beds must be sized to remove all H2O and provide for interference from other molecules in order to remove all H2S.
The adsorption process usually occurs at moderate pressure. Ionic bonds tend to achieve an optimum performance near 450 psig but can operate in a wide range of pressures.
Hydrogen sulfide can be selectively removed to meet 0.25 grain/100 scf (4ppm) specification. However, this reduction requires regeneration of the bed at 600°F (315°C) for extended time (Usually one hour or more depending upon process conditions) with the potential for COS formation if 4A is used (Molecular sieves are classified by their nominal pore diameter in Angstroms. See Chapter 5 for details.).
The problem of COS formation during processing according to the reaction in Equation 6-18, has been extensively studied.
H2S + CO2 ↔ COS + H2O Eq. 6-18
Molecular sieve products have been developed that do not catalyze COS formation. The central zone in the regeneration cycle is most favorable to COS formation.
Regeneration of a molecular sieve bed concentrates the H2S into a small regeneration stream which must be treated or disposed of.
Figure 6-12 shows a typical flow diagram for removal of H2S from natural gas. The configuration is similar to that for dehydration but with the significant difference that the regeneration gas contains high quantities of H2S as well as water as it leaves the adsorbent bed and, thus, must be treated. (Chapter 5 presents a thorough discussion of adsorption.) The flow configuration shows the first bed in the adsorption cycle, the second bed cooling down after regeneration, and the third bed undergoing regeneration with hot gas.
Fig. 6-12. Integrated Natural Gas Desulfurization Plant
Chi and Lee (1973) studied the coadsorption of H2S, CO2, and H2O on a 5A molecular sieve from a natural gas mixture under a variety of conditions.
Figure 6-13, from their paper, shows a typical concentration versus time curve.
In the figure, y is the concentration in the exit stream and yo is the concentration in the inlet to the bed. The gas that entered the bed was saturated with H2O and contained both CO2 (1.14 mol%) and H2S (0.073 mol%.). Because the CO2 content of the gas was 15.6 times that of the H2S, the bed quickly saturated with CO2, and its breakthrough was almost instantaneous. As the H2S was adsorbed and moved down the column, it displaced the CO2 and, consequently, after approximately 30 minutes, the CO2 exit concentration peaked at a value greater than its inlet concentration.
The same phenomenon occurs when the H2S is displaced by the water.
Because the H2S must be removed to extremely low levels, the bed is effectively exhausted from the perspective of H2S purification shortly after H2S breakthrough occurs and, thus, well before the bed is totally saturated.
Fig. 6.13 Effluent H2S and CO2 concentration from adsorption bed as a function of Time.
In general, the sieve bed can be designed to dehydrate and sweeten simultaneously.
The Engineering Data Book (GPSA) notes that a key point in adsorber design is to properly design for treatment of the regeneration gas because the peak H2S concentration may be 30 times the H2S concentration in the feed.
Care should be taken to minimize mechanical damage to the solid crystals as this will decrease the bed’s effectiveness. The main cause of mechanical degradation is the sudden pressure and/or temperature changes that may occur when switching from adsorption to regeneration cycles. Proper instrumentation can significantly extend bed life.
The molecular sieve process is limited to small gas streams operating at moderate pressures. It is generally used for polishing applications following one of the other processes.
6.7.9 Oxorbon
Oxorbon is an alternative solid bed material which consists of activated carbon impregnated with potassium iodide (KI). Donau Carbon markets such a carbon for the removal of H2S and mercaptans. The adsorbed H2S is converted to elemental sulfur by catalytic reaction under the presence of oxygen. The resulting sulfur is fixed on the pores of the activated carbon. Sulfur loadings as high as 60% of the carbon mass have been reported and sulfide concentrations below 1 ppmv are claimed.
6.8 Direct Conversion Processes (Liquid Redox)
Chemical and physical solvent processes remove acid gas from the natural gas stream but release H2S and CO2 when the solvent is regenerated. The release of H2S to the atmosphere is limited by environmental regulations. Acid gases could be routed to an incinerator/flare, which would convert the H2S to SO2. Environmental regulations restrict the amount of SO2 vented or flared. Direct conversion processes use chemical reactions to oxidize H2S and produce elemental sulfur. These processes are generally based either on the reaction of H2S and O2 or H2S and SO2. Both reactions yield water and elemental sulfur. These processes are licensed and involve specialized catalysts and/or solvents.
The redox agent is then regenerated by reaction with air in an oxidizer vessel.
Liquid oxidation−reduction (redox) processes use iron as an oxidizing agent in solution. The process involves four basic steps:
1. Removal of H2S from the gas by absorption into a caustic solution
2. Oxidation of the HS− ion to elemental sulfur via the oxidizing agent
3. Separation and removal of sulfur from the solution
4. Regeneration (i.e., oxidation) of the oxidizing agent by use of air
The processes share the general chemistry shown below.
• H2S absorption using an alkaline solution:
H2S (g) ↔ H2S (sol’n) Eq 6-19
H2S (sol’n) ↔ H+ + HS− Eq 6-20
Note that this is a non-selective, basic scrubbing solution; thus, additional acidic components (CO2 , HCN) of the sour gas will be soluble in the scrubbing solution. While CO2 is partially absorbed in the alkaline solution, it does not participate in the redox reactions. CO2 may however increase the rate of consumption of caustic or buffering compounds. The fate of other reduced sulfur species (COS, CS2, RSH, RSR) depends on the particular process being considered.
• Conversion to Elemental Sulfur:
HS− + H+ + 1/2 O2 (sol’n) → H2O (l) + S Eq 6-21
Note that in the absence of an auxiliary redox reagent (ARR) the reaction shown as Eq 21-24 is slow and nonspecific. The addition of an (ARR) increases the rate of reaction and directs the oxidation to elemental sulfur.
HS− + ARR(OX) → S + H+ + ARR (RED) Eq 6-22
where (OX) denotes the oxidized form and (RED) denotes the reduced form of the ARR.
• Regeneration of the Spent ARR Using Air:
O2 (g) ↔ O2 (sol’n) Eq 6-23
ARR (RED) + 1/2 O2 (sol’n) + 2 H+ → ARR(OX) + H2O(l) Eq 6-24
The overall simplified chemistry of liquid redox processes is thus:
H2S(g) + 1/2 O2 (g) → H2O(l) + S Eq 6-25
While the oxidizing agent forms the desired reaction, it also reacts with the sulfur species to form metal sulfides that precipitate from solution. To avoid the metal sulfide reaction, the solution contains chelating agents (organic compounds that bind with the metal ion to restrict its reactivity but still permit electron transfer for oxidation−reduction reactions). A common chelating agent for iron is EDTA (ethylenediaminetetraacetic acid). Several processes, including Lo-Cat II®, SulFerox ®, and Sulfint-HP, use iron and chelating agents.
6.8.1 Stretford Process
The Stretford process involves the use of vanadium salts as the ARR. The process has been extensively used in Europe. However environmental concerns around the discharge of vanadium compounds has limited its use.
Figure 6.14 shows a simplified diagram of the Stretford process. The gas stream is washed with an aqueous solution of sodium carbonate, sodium vanadate, and anthraquinone disulfonic. An oxidized solution is delivered from the pumping tank to the top of the absorber tower where it contacts the gas stream in a counter-current flow.
The bottom of the absorber tower consists of a reaction tank from which the reduced solution passes to the solution flash drum, which is situated above the oxidizer. The reduced solution passes from here into the base of the oxidizer vessel. Hydrocarbon gases, which have been dissolved in the solution at the plant pressure, are released from the top of the flash drum.
Air is blown into the oxidizer, and the main body of the solution, now reoxidized, passes into the pumping tank. The sulfur is carried to the top of the oxidizer by froth created by the aeration of the solution and passes into the thickener.
The function of the thickener is to increase the weight percent of sulfur that is pumped to one of the alternate sulfur recovery methods of filtration, filtration and autoclaves, centrifugation or centrifugation with heating.
Chemical reactions involved are:
H2S+Na2CO3 → NaHS+NaHCO3 Eq 6-26
Sodium carbonate provides the alkaline solution for initial adsorption of H2S and the formation of hydrosulfide (HS). The hydrosulfide is reduced in a reaction with sodium meta vanadate to precipitate sulfur
HS- +V+5 → S +V+4 Eq 6-27
Anthraquinone disulfonic acid (ADA) reacts with 4-valent vanadium and converts it back to 5 valent
V+4 +ADA → V+5 +ADA (reduced) Eq 6-28
Oxygen from the air converts the reduced ADA back to the oxidized state as shown below:
Reduced ADA+O2 → ADA+H2O Eq 6-29
The overall reaction is
2H2S+O2 → 2H2O+2S Eq 6-30
6.8.2 Lo-Cat process
This is a liquid-redox process. The LO-CAT® process is a patented, wet scrubbing, liquid redox system that uses a chelated iron solution to convert H2S to innocuous, elemental sulfur. It does not use any toxic chemicals and does not produce any hazardous waste byproducts. In general, the LO-CAT process can directly treat a gas stream, or treat the H2S containing stream from an Acid Gas Removal unit, although the direct treat capability is limited to low pressure streams.
Processes employ high iron concentration reduction-oxidation technology for the selective removal of H2S (not reactive to CO2) to <4 ppm in both low and high pressure gas streams. Figure 6-15 shows a simplified flow schematic of the LO-CAT process.
Acid gas stream is contacted with the solution where H2S reacts with, and reduces the chelated iron and produces elemental sulfur. The iron is then regenerated by bubbling air through the solution. Heat is not required for regeneration. The reactions involved are exothermic (give off heat):
Absorption/reduction: 2Fe3+ +H2S → 2Fe2+ +S+2H+ Eq 6-31
Regeneration/oxidation: 2Fe2+ +½ O2 +2H+ → 2Fe3+ +H2O Eq 6-32
Overall chemistry: H2S + ½ O2 → S + H2O Eq 6-33
Fig 6.14. Simplified flow schematic of the Stretford process.
Fig. 6.15. Simplified flow schematic of the LO-CAT process.
6.8.3 Sulferox process
SulFerox is a Shell proprietary Iron Redox process, whereby a sour gas stream, containing hydrogen sulphide, is contacted with a liquid, containing soluble ferric (Fe3+) ions. In the process the H2S is oxidized to elemental sulphur and the Fe3+ is reduced to ferrous (Fe2+) ions. The system is regenerable, Fe2+ is subsequently reconverted to Fe3+ by oxidation with air. Sulphur is recovered from the aqueous solution as a moist cake.
The process offers low capital and operating costs through the use of a high concentration iron chelate solution and effective control of chelate degradation. The process also offers a patented contactor design to improve the overall efficiency of the process which further reduces capital costs. The optimum application for SulFerox is in the one to twenty tons per day range of recovered sulfur.
Typical Sulferox process equipment is illustrated in Figure 6.16. In this process, the sour gas is contacted in a small contactor in co-current flow with a water solution containing about 4% ferric iron ions, held in solution by proprietary chelate ligand. The H2S ionizes in the solution, and the ferric ions exchange electrons with the sulfur ions to form ferrous ions and elemental sulfur. The gas and the solution leave the contactor and are flowed into a separator. The gas out of the separator is sweet and requires further treating for dewpoint control. The solution is flowed into additional vessels for separating the elemental sulfur and for restoring the ferrous iron ions back to the active ferric iron state by contacting the solution with air.
In general, the Sulferox process can directly treat a gas stream, or treat the H2S containing stream from an Acid Gas Removal unit, although the direct treat capability is limited to low pressure streams.
Fig. 6.16. Simplified flow schematic of the sulferox process.
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Fundamentals of Oil and Gas Processing
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Basics of Gas Field Processing
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