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Gas Sweetening and Sulfur Recovery - Chapter 6 - Part 2 B

6.8.4 IFP Process
The IFP process was developed by the Institut Francais du Petrole. The process reacts H2S with SO2 to produce water and sulfur.
The overall reaction is H2S+SO2 → H2O+2S
Figure 6.17 shows a simplified diagram of the IFP process. The process involves mixing the H2S and SO2 gases and then contacting them with a liquid catalyst in a packed tower. Elemental sulfur is recovered in the bottom of the tower. A portion of this must be burned to produce the SO2 required to remove the H2S.
The most important variable is the ratio of H2S to SO2 in the feed. The ratio is controlled by analyzer equipment to maintain the system performance.


Fig. 6.17 .Simplified flow schematic of the IFP process.

6.9 Distillation Process
Distillation processes uses cryogenic distillation to remove acid gases from a gas stream. The process is applied to remove CO2 for LPG separation or where it is desired to produce CO2 at high pressure for reservoir injection or other use.
The process consists of two, three, or four fractionating columns. The gas stream is first dehydrated and then cooled with refrigeration and/or pressure reduction.

6.9.1 Three-Column System
The three-column system is used for gas streams containing <50% CO2. The first column operates at 450-650 psig and separates a high-quality methane product in the overhead. Temperatures in the overhead are from 0 to -140 0F (-18 to -95 0C). The second column operates at a slightly lower pressure and produces a CO2 stream overhead, which contains small amounts of H2S and methane. The bottom product contains H2S and the ethane plus components. The third column produces NGL liquids, which are recycled back to the first two columns.
NGL liquids recycled prevent CO2 solid formation in the first column and aid in the breaking of the ethane/CO2 azeotrope in the second column to permit high ethane recoveries.
6.9.2 Four-Column System
The four-column system is used where CO2 feed concentration exceeds 50%. The initial column in this scheme is a de-ethanizer. The overhead product, a CO2/methane binary, is sent to a bulk CO2 removal column and de-methanizer combination. CO2 is produced as a liquid and is pumped to injection or sales pressure.

6.9.3 Two-Column System
The two-column system is used when a methane product is not required and is thus produced with the CO2. Very high propane recoveries may be achieved; however, little ethane recovery is achieved. These processes require feed gas preparation in the form of compression and dehydration, which adds to their cost. Such systems are finding applications in enhanced oil recovery (EOR) projects.
6.10 Sulfur Recovery (The Claus Process)
Sulfur is present in natural gas principally as hydrogen sulfide (H2S) and, in other fossil fuels, as sulfur-containing compounds which are converted to hydrogen sulfide during processing. The H2S, together with some or all of any carbon dioxide (CO2) present, is removed from the natural gas or refinery gas by means of one of the gas treating processes pre-described. The resulting H2S-containing acid gas stream is flared, incinerated, or fed to a sulfur recovery unit.
The Claus process is used to treat gas streams containing high (above 50%) concentrations of H2S.
The Claus process as used today is a modification of a process first used in 1883 in which H2S was reacted over a catalyst with air (oxygen) to form elemental sulfur and water.
H2S+1/2 O2 → S +H2O Eq. 6-34

Control of this highly exothermic reaction was difficult and sulfur recovery efficiencies were low. In order to overcome these process deficiencies, a modification of the Claus process was developed and introduced in 1936 in which the overall reaction was separated into:

1- A highly exothermic thermal or combustion reaction section in which most of the overall heat of reaction (from burning one-third of the H2S and essentially 100% of any hydrocarbons and other combustibles in the feed) is released and removed, and
2- A moderately exothermic catalytic reaction section in which sulfur dioxide (SO2) formed in the combustion section (step 1) reacts with unburned H2S to form elemental sulfur. The principal reactions taking place (neglecting those of the hydrocarbons and other combustibles) can then be written as follows:

Thermal or Combustion Reaction Section
H2S+3/2 O2 → SO2 +H2O Eq. 6-35

Combustion and Catalytic Reaction Sections
SO2 +2H2S → 3S+ 2H2O Eq. 6-36

Overall Reaction
3H2S+3/2 O2 → 3S+ 3H2O Eq. 6-37

Figure 6-18 shows a simplified flow diagram of a two-stage Claus process plant. The first stage of the process converts H2S to sulfur dioxide and to sulfur by burning the acid gas stream with air in the reaction furnace. This provides SO2 for the next phase of the reaction.
Gases leaving the furnace are cooled to separate out elemental sulfur formed in the thermal stage. Reheating, catalytically reacting, and sulfur condensation removes additional sulfur. Multiple reactors are provided to achieve a more complete conversion of the H2S. Condensers are provided after each reactor to condense the sulfur vapor and separate it from the main stream.
Conversion efficiencies of 94-95% can be attained with two catalytic stages while up to 97% conversion can be attained with three catalytic stages.
The efficiencies are dictated by environmental concerns; the effluent gas (SO2) is either vented, incinerated, or sent to a “tail gas treating unit.”

Fig. 6-18. Simplified process flow schematic for a two-stage Claus process plant.

For the usual Claus plant feed gas composition (water-saturated with 30-80 mol % H2S, 0.5-1.5 mol % hydrocarbons, the remainder CO2), the modified Claus process arrangement results in thermal section (burner) temperatures of about 1800 to 2500°F.
Sulfur recovery would be expected to be lower for a feed gas from a refinery than for a wellhead treater because of higher hydrocarbon content.
Conversion of H2S to elemental sulfur is favored in the reaction furnace by higher operating temperatures of 1800°F and in the catalytic converters by lower operating temperatures of less than 700°F.
To attain an overall sulfur recovery level above about 70%, the thermal, or combustion, section of the plant is followed by one or more catalytic reaction stages. Sulfur is condensed and separated from the process gases after the combustion section and after each catalytic reaction stage in order to improve equilibrium conversion. The process gases must be reheated prior to being fed to the catalytic reaction stage in order to maintain acceptable reaction rates and to ensure that the process gases remain above the sulfur dewpoint as additional sulfur is formed. Figure 6-19 is the flow sheet of an example three-stage Claus sulfur recovery plant; Figure 6-20 shows the mechanical arrangement of an example small, package-type, two-stage Claus plant.
Gases leaving the final sulfur condensation and separation stage may require further processing. These requirements are established by local, or national regulatory agencies.
These requirements can be affected by the size of the sulfur recovery plant, the H2S content of the plant feed gas, and the geographical location of the plant.



Fig. 6-19. Three-Stage Sulfur Plant. (Straight-Through Operating with Acid Gas-Fueled Inline Burners for Reheating)


Fig. 6-20. Example Package-Type Sulfur Plant

6.10.1 Claus Process Considerations
The Claus sulfur recovery process includes the following process operations:
• Combustion — burn hydrocarbons and other combustibles and 1/3 of the H2S in the feed.
• Waste Heat Recovery — cool combustion products. Because most Claus plants produce 150-500 psig steam (365-470°F), the temperature of the cooled process gas stream is usually about 600-700°F.
• Sulfur Condensing — cool outlet streams from waste heat recovery unit and from catalytic converters. Low temperature of the cooled gas stream is usually about 350°F or 260-300°F for the last condenser.
• Reheating — Reheat process stream, after sulfur condensation and separation, to a temperature high enough to remain sufficiently above the sulfur dewpoint, and generally, for the first converter, high enough to promote hydrolysis of COS and CS2 to H2S and CO2.
COS + H2O → CO2 + H2S Eq 6-38
CS2 + 2H2O → CO2 +2H2S Eq 6-39
• Catalytic Conversion — Promote reaction of H2S and SO2 to form elemental sulfur.

6.10.2 Process Variations
Several variations of the basic Claus process have been developed to handle a wide range of feed gas compositions. Some of these are shown in Figure 6-21. Straight-through operation results in the highest overall sulfur recovery efficiency and is chosen whenever feasible.
Table. 6-12 can be used as a guide in Claus process selection.

Fig. 6-21. Claus Process Variations


Table. 6-12. Claus Plant Configurations
6.10.3 Combustion Operation
Most Claus plants operate in the "straight-through" mode.
The combustion is carried out in a reducing atmosphere with only enough air (1) to oxidize one-third of the H2S to SO2, (2) to burn hydrocarbons and mercaptans, and (3) for many refinery
Claus units, to oxidize ammonia and cyanides. Air is supplied by a blower and the combustion is carried out at 3-14 psig, depending on the number of converters and whether a tail gas unit is installed downstream of the Claus plant.
Numerous side reactions can also take place during the combustion operation, resulting in such products as hydrogen (H2), carbon monoxide (CO), carbonyl sulfide (COS), and carbon disulfide
(CS2). Thermal decomposition of H2S appears to be the most likely source of hydrogen since the concentration of H2 in the product gas is roughly proportional to the concentration of H2S in the feed gas. Formation of CO, COS, and CS2 is related to the amounts of CO2 and/or hydrocarbons present in the feed gas.
Heavy hydrocarbons, ammonia, and cyanides are difficult to burn completely in a reducing atmosphere. Heavy hydrocarbons may burn partially and form carbon which can cause deactivation of the Claus catalyst and the production of off color sulfur. Ammonia and cyanides can burn to form nitric oxide (NO) which catalyzes the oxidation of sulfur dioxide (SO2) to sulfur trioxide (SO3); SO3 causes sulfation of the catalyst and can also cause severe corrosion in cooler parts of the unit. Unburned ammonia may form ammonium salts which can plug the catalytic converters, sulfur condensers, liquid sulfur drain legs, etc. Feed streams containing ammonia and cyanides are sometimes handled in a special two-combustion stage burner or in a separate burner to ensure satisfactory combustion.
Flame stability can be a problem with low H2S content feeds (a flame temperature of about 1800°F appears to be the minimum for stable operation).

The split flow, sulfur recycle, or direct oxidation process variations often are utilized to handle these H2S-lean feeds; but in these process schemes, any hydrocarbons, ammonia, cyanides, etc. in all or part of the feed gas are fed unburned to the first catalytic converter. This can result in the cracking of heavy hydrocarbons to form carbon or carbonaceous deposits and the formation of ammonium salts, resulting in deactivation of the catalyst and/or plugging of equipment.

A method of avoiding these problems while still improving flame stability is to preheat the combustion air and/or acid gas, and to operate "straight-through". An example of such an arrangement is shown in Figure 6-22. Steam-, hot oil-, or hot gas-heated exchangers and direct fired heaters have been used. The air and acid gas are usually heated to about 450-500°F. Sometimes split flow is combined with acid-gas preheat. Other methods of improving flame stability are to use a high intensity burner, to add fuel gas to the feed gas, or to use oxygen or oxygen-enriched air for combustion.

6.10.4 Claus Unit Tail Gas Handling
The tail gas from a Claus unit contains N2, CO2, H2O, CO, H2, unreacted H2S and SO2, COS, CS2, sulfur vapor, and entrained liquid sulfur. Because of equilibrium limitations and other sulfur losses, overall sulfur recovery efficiency in a Claus unit usually does not exceed 96-97%. Venting of this tail gas stream without further processing is seldom permitted; the minimum requirement is normally incineration, the principal purposes of which are to reduce H2S concentrations to a low level (which value will depend on the local regulations) and to provide the thermal lift for dispersion of SO2 upon release to atmosphere through a stack. Depending upon the size of the
Claus unit, the H2S content of the feed gas, and the geographical location, a tail gas cleanup process may be required in order to reduce emissions to the atmosphere.



Fig. 6-22. Sulfur Recovery Process with Acid Gas and Air Preheat

Incineration
Incineration of the H2S (as well as the other forms of sulfur) in the Claus plant tail gas to SO2 can be done thermally or catalytically. Thermal oxidation normally is carried out at temperatures between 900°F and 1500°F in the presence of excess oxygen. Most thermal incinerators are natural draft operating at sub-atmospheric pressure with air flow controlled with dampeners; the excess oxygen level varies between 20% and 100%. A typical concentration of oxygen in the stack effluent is 2.0%. Although the Claus unit tail gas contains some combustibles — for example, H2S, COS, CO, CS2, H2, and elemental sulfur (in the case of "split-flow" plants, some hydrocarbons) — these combustibles are at too low a concentration to burn since they generally amount to less than 3% of the total tail gas stream. The entire tail gas stream must therefore be incinerated at a high enough temperature for oxidation of sulfur and sulfur compounds to SO2.
Incinerator fuel consumption can be reduced significantly by utilizing catalytic incineration. This involves heating the tail gas stream to about 600-800°F with fuel gas and then passing the heated gas along with a controlled amount of air through a catalyst bed. Catalytic incinerators are normally forced draft, operating at a positive pressure in order to maintain closer control of excess air. Catalytic incineration is a proprietary process which should be considered where fuel costs for conventional (thermal) incineration are high.
Another method of improving overall fuel economy involves recovering heat from the incinerator outlet gases. Saturated steam at pressures ranging between 50 psig and 450 psig has been produced, and saturated steam has been superheated, using waste heat from the incinerator outlet gases.
Incinerators with waste heat recovery are normally forced draft operating at a positive pressure.
Fuel required for thermal incineration is determined by the amount of heat needed to heat the Claus tail gas, air, and fuel to the required temperature. Normally the incinerator is sized for at least 0.5 second residence time, and sometimes for as much as 1.5 seconds residence time. Generally, the longer the residence time, the lower the incinerator temperature needed to meet the environmental requirements. This is illustrated by Figure 6-23 which shows the relationship between residence time and temperature for a typical installation to meet a maximum H2S requirement of 10 ppmv.


Fig. 6-23. Typical Relationship Between Incinerator Residence Time and Required Temperature.

The incinerator and stack can sometimes be combined into a single vessel. The incinerator is the

Tail Gas Clean-up Processes (TGCU)
All of the Claus tail gas cleanup (TGCU) processes fit roughly into four categories:
Processes based primarily on the continuation of the Claus reaction to produce additional sulfur under more favorable equilibrium conditions than normally found in the Claus units, either through operation at temperatures below the sulfur dewpoint or in the liquid phase at a temperature above the melting point of sulfur.
Processes based on converting all the sulfur components in the tail gas to SO2 and recovering the SO2 for further processing.
Processes based on converting all the sulfur in the Claus unit tail gas to H2S, then recovering sulfur from this H2S.
Processes that directly oxidize the tail-gas H2S to sulphur.


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Operational Aspects
Overall Claus plant conversion efficiency is maximized by maintaining the stoichiometric H2S:SO2 ratio of 2:1 in the process gas to the catalytic converters. The most suitable point for this determination is at the outlet of the last sulfur condenser because a slight change in the air:acid gas ratio at the front of the plant will result in a significant change in the H2S:SO2 ratio in the tail gas and in the theoretical overall sulfur recovery. An H2S:SO2 ratio in the tail gas of between 1:1 and 3:1 can be considered normal although the desired goal should be a 2:1 ratio.
Because of the effect of temperature upon the Claus reaction equilibrium, control of temperatures at various points in the process sequence is important. Unexpected changes in operating temperatures usually denote changes in conversion efficiency. For example, a decrease in the temperature rise across a catalytic converter bed is an indication of declining catalyst activity which may be caused by adsorption of elemental sulfur on the active surface area of the catalyst. Operating the catalyst bed at a temperature 50-100°F higher than normal for 24-48 hours will remove this sulfur from the catalyst and can restore its activity. 
6.11 Gas Permeation Process (Membranes)
Membranes are thin semipermeable barriers that selectively separate some compounds from others.
Membranes are used in natural gas processing for dehydration, fuel-gas conditioning, and bulk CO2 removal, but presently CO2 removal is by far the most important application. In some applications, membranes are used to recover CO2 from EOR floods for recycle injection into oil and gas reservoir.

6.11.1 Membrane Fundamentals
Membranes do not act as filters where small molecules are separated from larger ones through a medium of pores. They operate on the principle of solution-diffusion through a nonporous membrane. Highly solubilized components dissolve and diffuse through the membrane.

Relative permeation rates
• Most soluble (fastest gases)
H2O, H2, H2S, CO2, O2
• Least soluble (slowest gases)
N2, CH4, C2+
CO2 first dissolves into the membrane and then diffuses through it. Membranes allow selective removal of fast gases from slow gases.
Membranes do not separate on the basis of molecular size. Separation is based on how well different compounds dissolve into the membrane and then diffuse it.
Fick’s law (known as Basic Flux Equation) is used to approximate the solution-diffusion process. It is expressed as
Ji (Si Di pi)/L Eq. 6-40
where
J is the flux of component i, that is, the molar flow of component i through the membrane per unit area of membrane,
Si is the solubility term,
Di is the diffusion coefficient,
pi is the partial pressure difference across the membrane, and
L is the thickness of the membrane.
Customarily, Si, and Di are combined into a single term, the permeability, Pi, and thus divides Fick’s law into two parts:
Pi /L, which is membrane dependent and
pi, which is process dependent.
(Note that Pi/L is not only dependent on the membrane but also dependent on operating conditions, because Si and Di depend on both temperature and pressure. Pi also depends weakly upon the composition of the gases present).
All the mixture components have a finite permeability, and the separation is based upon differences in them. Customarily, selectivity, 1-2, is used, which is the ratio of two permeabilities, P1/P2, a term important in process design and evaluation. An  of 20 for CO2/CH4 means that CO2 moves through the membrane 20 times faster than does methane.

6.11.2 Membrane Selection Parameters
Permeability:
High permeability results in less membrane area required for a given separation and a lower system cost.
Selectivity:
High selectivity results in lower losses of hydrocarbons as CO2 is removed and a higher volume of salable product.
Unfortunately, high CO2 permeability does not correspond to high selectivity. A choice must be made between a highly selective, or permeable, membrane and somewhere between on both parameters. The usual choice is to use a highly selective material and then make it as thin as possible to increase the permeability. Reduced thickness makes the membrane extremely fragile and therefore unusable.
In the past, membrane systems were not a viable process because the membrane thickness required to provide the mechanical strength was so high that the permeability was minimal.

6.11.3 Membrane Structure Types
Asymmetric Membrane Structure
An asymmetric membrane structure features a single polymer consisting of an extremely thin nonporous layer mounted on a much thicker and highly porous layer of the same material, as opposed to a homogenous structure, where membrane porosity is more-or-less uniform throughout. Figure 6-24 is an example of an asymmetric membrane.
Nonporous layer
Meets the requirements of the ideal membrane, that is, highly selective, and thin.
Porous layer
Provides mechanical support and allows the free flow of compounds that permeate through the nonporous layer.

Fig. 6-24. Asymmetric membrane structure, and a Composite membrane structure.

Composite Membrane Structure
The disadvantages of the asymmetric membrane structure are they are composed of a single polymer; they are expensive to make out of exotic, highly customized polymers; and they are produced in small quantities.
These drawbacks are overcome by producing a composite membrane.
The composite membrane consists of a thin selective layer made of one polymer mounted on an asymmetric membrane, which is made of another polymer.
The composite structure allows manufacturers to use readily available materials for the asymmetric portion of the membrane and specially developed polymers, which are highly optimized for the required separation and the selective layer.
Composite structures are being used in most newer advanced CO2 removal membranes because the proprieties of the selective layer can be adjusted readily without significantly increasing membrane cost.

6.11.4 Carbon Dioxide Removal from Natural Gas
Many different types of membranes have been developed or are under development for industrial separations, but for CO2 removal, the industry standard is presently cellulose acetate. In these membranes are of the solution-diffusion type, in which a thin layer (0.1 to 0.5 μm) of cellulose acetate is on top of a thicker layer of a porous support material. Permeable compounds dissolve into the membrane, diffuse across it, and then travel through the inactive support material. The membranes are thin to maximize mass transfer and, thus, minimize surface area and cost, so the support layer is necessary to provide the needed mechanical strength.

6.11.5 Membrane Elements
Commercial membrane configurations are either hollow fiber elements or flat sheets wrapped into spirally wound elements. Presently, about 80% of gas separation membranes are formed into hollow fiber modules, like those shown in Figures 6-26 & 6.27.

Flat Sheet (Spiral Wound)
In the spiral wound element shown in Figure 6-25, two membrane sheets are separated by a permeate spacer and glued shut at three ends to form an envelope or leaf. Many of these leaves, separated by feed spacers, are wrapped around the permeate tube, with the open end of the leaves facing the tube. Feed gas travels along the feed spacers, the permeating species diffuse through the membranes and down the permeate spacers into the permeate tube, and the residue gas exits at the end. The gas flow is cross flow in this configuration.
The spiral configuration is inherently more resistant than the hollow fiber membranes to trace components that would alter the polymer permeability. It also allows a wider range of membrane materials to be used. However, the hollow fiber membranes are cheaper to fabricate, and thus dominate the field.
Optimization involves the number of envelopes and element diameter.
Number of envelopes
The permeate gas must travel the length of each envelope. Having many shorter envelopes makes more sense than having a few longer ones because pressure drop is greatly reduced in the former case.
Element diameter
A larger bundle diameter allow better packing densities but increases the element tube size and decreases cost. A larger diameter also increases the element weight, which makes the elements more difficult to handle during installation and replacement.

Hollow Fiber
As shown in Figures 6-27 and 6-27, very fine hollow fibers are wrapped around a central tube in a highly dense pattern. Feed gas flows over and between the fibers and some components permeate into them.
Permeate gas travels within the fibers until it reaches the permeate pot, where it mixes with the permeate from other fibers. The total permeate exits the element through a permeate pipe. Gas that does not permeate eventually reaches the element’s center tube, which is perforated. In this case, the central tube is for residual collection, not permeate collection.
The low-pressure, bore-feed configuration is a countercurrent flow configuration similar to a shell-tube heat exchanger with the gas entering on the tube side. It has the advantage of being more resistant to fouling because the inlet gas flows through the inside of the hollow fibers.
However, the mechanical strength of the membrane limits the pressure drop across the membrane. The configuration is only used in low-pressure applications, such as air separation and air dehydration.
To handle high pressures, the permeate flows into the hollow fiber from the shell side. This feature makes the membrane much more susceptible to plugging, and gas pretreatment is usually required. The gas flow is cross current and provides good feed distribution in the module. This configuration is widely used to remove CO2 from natural gas.

Fig. 6-25 Spiral wound membrane element. (UOP - LLC)

Spiral Wound Versus Hollow Fiber
Spiral Wound Hollow fiber
• Installed in horizontal vessels
• Operate at higher allowable operating pressures 1085 psig (75 barg) and thus have higher driving force available for permeation
• More resistant to fouling
• Have a long history of service in natural gas sweetening
• Perform best with colder inlet stream gas temperatures
• Do not handle varying inlet feed quality as well as hollow fiber units installed in vertical vessels
• Require extensive pretreatment equipment with high inlet stream liquid hydrocarbon loading • Characteristics of hollow fiber membranes
• Installed in vertical vessels
• Offer a higher packing density
• Operate at lower inlet stream pressures 580 psig (40 barg)
• Handle higher inlet stream hydrocarbon loading better than spiral wound units
• Require inlet feed gas chilling
• Hollow fiber based plants are typically smaller than spiral wound-based plants
• Handle varying inlet feed quality better than spiral wound units installed in horizontal vessels.
Finer fibers give higher packing density, but larger fibers have lower permeate pressure drops and so they use the pressure driving force more efficiently.
Table. 6-13. Spiral Wound Versus Hollow Fiber.
Membrane Modules
Once the membranes have been manufactured into elements, they are joined together and inserted into a tube (Figure 6-28). Multiple tubes are mounted on skids in either a horizontal or vertical orientation, depending on the membrane company.

Fig. 6-26. Hollow-fiber membrane element.

Fig. 6.27 Cutaway view of the two module configurations used with hollow fiber membranes.


Fig. 6-28. Cutaway view of spiral wound membrane module. (UOP - LLC.)


6.11.6 Membrane Design Considerations
Process Variables Affecting Design are:

6.11.6.1 Flow Rate
A maximum acceptable feed gas rate per unit area applies to the membrane, and required membrane area is directly proportional to the flow rate. Membrane units perform well at reduced feed rates, but their performance drops when design flow rates are exceeded. Additional modules are added in parallel to accept higher flow rates.
The percentage of hydrocarbon losses (hydrocarbon losses/feed hydrocarbons) remains the same at different flow rates.

6.11.6.2 Operating Temperature
Increased operating temperature increases permeability but decreases selectivity.
Membrane area requirement is decreased, but the hydrocarbon losses and recycle compressor power for multistage systems are increased (Figure 6-29).
Because membranes are organic polymers, they have a maximum operating temperature that depends upon the polymer used. Exceeding this temperature will degrade membrane material and shorten the useful life of the unit.

6.11.6.3 Feed Pressure
An increase in feed pressure decreases both membrane permeability and selectivity, but at the same time creates a greater driving force across the membrane that results in a net increase in permeation through the membrane and a decrease in the membrane area requirements (Figure 6-30). Increasing the maximum operating pressure results in a less expensive and smaller system. Limiting factors are the maximum pressure limit for the membrane elements and the cost and weight of equipment at the higher pressure rating.

Fig. 6-29. Effect of operating temperature.

Fig. 6-30. Effect of feed pressure.

6.11.6.4 Permeate Pressure
Exhibits the opposite effects of feed pressure Lowers the permeate pressure Increases the driving force, and lowers the membrane area requirements. Unlike feed pressure, permeate pressure has a strong effect on hydrocarbon losses (Figure 6-31).
Pressure difference across the membrane is not the only consideration. Pressure ratio across the membrane is strongly affected by the permeate pressure.
For example, a feed pressure of 1305 psig (90 bar) and a permeate pressure of 43.5 psig (3 bar) produce a pressure ratio of 30. Decreasing the permeate pressure to 14.5 psig (1 bar) increases the pressure ratio to 90 and has a dramatic effect on system performance.


Fig. 6-31. Effect of permeate pressure.

Desirable to achieve the lowest possible permeate pressure Important consideration when deciding how to further process the permeate stream.
For example, if permeate stream must be flared, then the flare design must be optimized for low pressure drop. If permeate stream must be compressed to feed the second membrane stage or injected into a well, the increased compressor horsepower and size at lower permeate pressure must be balanced against the reduced membrane area requirements.

6.11.6.5 CO2 Removal
For a constant sales gas CO2 specification, an increase in feed CO2 increases membrane area requirement and increases hydrocarbon losses (more CO2 must permeate, and so more hydrocarbons permeate). This is shown in Figure 6-32.


Fig. 6-32. Effect of CO2 removal.

Membrane area requirement is determined by the percentage of CO2 removal rather than the feed or sales gas CO2 specifications themselves.
For example, a system for reducing a feed CO2 content from 10% to 5% is similar in size to one reducing the feed from 50% to 25%, or one reducing a feed from 1% to 0.5%, if all have a CO2 removal requirement of about 50%.
Traditional solvent or absorbent-based CO2 technologies have the opposite limitation.
Their size is driven by the absolute amount of CO2 that must be removed. For example, a system for CO2 removal from 50% to 25% is substantially larger than one reducing CO2 from 1% to 0.5%. For this reason, using membranes for bulk CO2 removal and using traditional technologies for meeting low CO2 specifications makes a lot of sense. Depending on the application, either one or both of the technologies could be used.
An increase in CO2 content in feed gas of an existing membrane plant will results in sales gas with higher CO2 content. An additional membrane area can be installed to meet the sales gas CO2 content, although with increased hydrocarbon losses. For example, if heater capacity is available, the membranes can be operated at a higher temperature to also increase capacity.
If an existing non-membrane system must be de-bottlenecked, installing a bulk CO2 removal system upstream of it makes good sense.

6.11.6.6 Environmental Regulations
Environmental regulations dictate what can be done with the permeate gas, specifically whether it can be vented (cold or hot vent) to the atmosphere or flared either directly or catalytically. Ninety five to ninety nine percent CO2 yields low Btu/scf content (flare requires a minimum of 250 Btu/scf to burn).

6.11.6.7 Location
Location often dictates a number of other issues, such as space and weight restrictions, level of automation, level of spares that should be available, and single versus multistage operation.
Fuel requirements can be obtained upstream of the membrane system, downstream of the pretreatment system, downstream of the membrane, and from the recycle loop in multistage systems.

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Gas Sweetening and Sulfur Recovery - Chapter 6 - Part 3 A

6.11.7 Operating Considerations

6.11.7.1 Flow Patterns
A number of different flow patterns can be used with membranes. Figures 6-33 and 6-34 show simplified examples for the CH4/CO2 separation. In the single-stage unit, the overall methane recovery is only 90.2%, but the process requires flow through only one membrane, and no recompression is needed. To increase methane recovery to 98.7%, a two-stage unit requires recompression of the first-stage permeate.
In the two-stage membrane processes, the first-stage permeate is in a second membrane stage. The permeate from the second stage, which has typically twice the CO2 content as the first stage permeate, is vented. The residue is either recycled or combined with the feed gas. A compressor is required to repressurize the first stage permeate before it is processed in the second stage. Two-stage designs provide higher hydrocarbon recoveries than two-step or one-stage designs, but they also require more compressor horsepower because more gas must be compressed.
Greater levels of methane recovery are obviously possible by application of three or more stages, but additional elements can quickly become uneconomical because of both membrane cost and recompression energy required.
The second law of thermodynamics dictates that energy is required for a separation. For membrane processes, this law translates into loss of pressure. However, in many processes, the cost of recompression, if needed, still makes membranes an attractive separation process. Many factors must be considered when deciding whether to use a single stage or multistage system. An economic analysis must be done to ensure that the cost of installing and operating a recycle compressor does not exceed the savings in hydrocarbon recovery.
Figure 6.35 plots the percentage hydrocarbon recovery versus percentage CO2 removal for one- and two-stage systems at certain process conditions.



Table 6.14. Typical membrane feed and permeate gas analysis


Percentage hydrocarbon recovery is defined as the percentage of hydrocarbon recovered to the sales gas versus the hydrocarbons in the feed gas. Hydrocarbon recovery of a two-stage is significantly better than that for a single-stage system. When deciding whether to use a single- or multistage approach, one must also consider the impact of the recycle compressor. Other considerations are additional hydrocarbons used as fuel, which increases the overall hydrocarbon losses, and the significant capital cost of the compressor and maintenance. For moderate CO2 removal applications (<50%), single-stage membrane systems usually provide better economic returns than do multistage systems.


Fig 6-33. A single-stage CO2/CH4 membrane separation process.
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Fig. 6-34. A two stage CO2/CH4 membrane separation process.

Fig. 6-35 Effect of number of stages.
6.11.8 Feed Gas Pretreatment
Because membranes are susceptible to degradation from impurities, pretreatment is usually required. The impurities possibly present in natural gas that may cause damage to the membrane include:
• Liquids. The liquids may be entrained in the feed to the unit or formed by condensation within the unit. Liquids can cause the membrane to swell, which results in decreased flux rates and possible membrane damage. Liquids can form internally by two mechanisms:
(1) Because of condensation of higher molar mass compounds caused by the cooling that occurs (Joule-Thomson effect) as the gas expands to a lower pressure through the membrane.
(2) Because CO2 and the lighter hydrocarbons diffuse more quickly than the heavier hydrocarbons, the dew point of the non-diffusing gas may increase to the point where condensation occurs.
• High-molar-mass hydrocarbons (C15+) such as compressor lube oils. These compounds coat the membrane surface and result in a loss of performance. The concentrations are low but the effect is cumulative.
• Particulates. These materials block the small flow passages in the membrane element. Blockage is lower for spiral-wound than for hollow-fiber elements (low flow area). Long-term particle flow into any membrane could eventually block it. Erosion of the membrane could also be a problem.
• Corrosion inhibitors and well additives. Certain of these compounds are destructive to membrane material.

Fig 6.36 Schematic of membrane-pretreating equipment (Traditional Pretreatment).

6.11.8.1 Traditional Pretreatment
A common method for pretreating the feed gas to a membrane system is shown in Figure 6-36. The coalescing filter removes any entrained liquids; the adsorbent bed takes out trace contaminants such as volatile organic compounds (VOC); the particulate filter removes any dust from the adsorbent bed; and the heater superheats the gas to prevent liquid formation in the membrane unit. The system shown has the following disadvantages:
• The adsorbent bed is the only unit that removes heavy hydrocarbons. Consequently, if the gas contains more heavy hydrocarbons than anticipated, or in the event of a surge of these materials, the adsorbent bed may become saturated in a relatively short time, and thus allow heavy hydrocarbons to contact the membrane.
• Only the heater provides superheat, and, consequently, if this unit fails, the entire membrane system must be shut down.
Other pretreatment methods that address the disadvantages discussed above.
Chiller
A chiller may be included to reduce the dew point of the gas and the heavy hydrocarbon content. Because chilling does not completely remove all heavy hydrocarbons, an adsorbent guard bed is still required. If deep chilling is necessary, steps must be taken to prevent hydrates from forming, either by dehydrating the gas upstream or by adding hydrate inhibitors. If inhibitors are added, they may need to be removed downstream of the chiller because some inhibitors may damage the membrane.
Turbo-expander
The turbo-expander serves the same purpose as a chiller, but has the benefit of being a dry system. It is smaller and lighter than the refrigeration system. A disadvantage is the net pressure loss, which must be taken up by the export compressor.
Glycol Unit
The glycol unit is added upstream of the chiller to prevent hydrate formation or freeze-up. An adsorbent guard bed is still required to remove heavy hydrocarbons but must be larger than it would normally be because it must also remove glycol carried over from the adsorber vessels.
6.11.8.2 Enhanced Pretreatment
It is common for an initial design, based on an extended gas analysis, to differ from actual analysis after the membrane system has been started up. For example, feed gas may be heavier than originally anticipated. Pretreatment systems may not have sufficient flexibility to handle a wide departure from design. Adsorbent beds may become fully saturated within a short time, leading to performance degradation. Preheaters may not be large enough to achieve feed temperatures that are much higher than designed.
A standard way to handle a gas that is heavier than expected is to operate the membranes at a higher temperature. Temperature increase increases the margin between gas dew point and operating temperature and thus prevents condensation in the membrane.
Figure 6.37 shows an enhanced pretreatment scheme that is more suitable for cases where one or more of the following is expected:
A wide variation in feed gas content
A significant amount of heavy hydrocarbons or other contaminants; or
A feed gas that may be heavier than analyzed, based on the known information from nearby wells or other locations.

Feed gas is first cooled down in a heat recovery exchanger, and any condensate formed is removed in a separator and a coalescer. Liquid-free gas then enters a regenerable adsorbent guard bed system where heavy hydrocarbons and other harmful components are completely removed. Water is removed along with the heavy hydrocarbons, and thus no upstream dehydration is required. The contaminant-free gas passes through a particle filter leaving the adsorbent guard bed system. Sometimes the product gas is cooled down in a chiller whose main purpose is to reduce the hydrocarbon dew point of the feed gas. Any condensate formed in the chiller is removed in a separator and the separator-outlet gas is routed to the feed cross exchanger.
Here, the gas cools down the system feed gas and obtains necessary superheat. Further superheat and control of membrane feed temperature are provided by a preheater.
Benefits of enhanced pretreatment are as follow:
Complete removal of heavy hydrocarbons
Unlike other pretreatment schemes, the absolute cutoff of heavy hydrocarbons is possible.
Regenerative system
Because adsorbent guard beds are regenerable, it is better able to handle fluctuations in the heavy hydrocarbon content of the feed gas than in traditional guard beds, which require frequent replacement of adsorbent material.
Ability to cope with varying feed composition
Cycle time can be adjusted to provide efficient treatment of a wide variety of feed compositions and heavy hydrocarbon contents.
Reliability
A system can be designed to operate satisfactorily even if one of its vessels is taken offline. Critical items in the pretreatment system are usually spared so they can be serviced or maintained without shutting the system down.
Efficiency
A system is able to provide a number of functions, such as removal of water, heavy hydrocarbons, and mercury, that would normally be provided by separate pieces of equipment. Heat recovery is implemented in the pretreatment scheme as well as within the system itself.

Fig. 6.37. Enhanced pretreatment flow scheme.


Operating Issues
Amines Membranes
User Comfort Level Very familiar Still considered new technology
Hydrocarbon Losses Very low Losses depend upon conditions
Meets Low CO2 Spec. Yes (ppm levels) No (<2% economics are challenging)
Meets Low H2S Spec. Yes (<4 ppm) Sometimes
Energy Consumption Moderate to high Low, unless compression used
Operating Cost Moderate Low to moderate
Maintenance Cost Low to moderate Low, unless compression used
Ease of Operation Relatively complex Relatively simple
Environmental Impact Moderate Low
Dehydration Product gas saturated Product gas dehydrated
Capital Cost Issues
Amines Membranes
Delivery Time Long for large systems Modular construction is faster
On-Site Installation Time Long Short for skid-mounted equipment
Pretreatment Costs Low Low to moderate
Recycle Compression Not used Use depends upon conditions
Table. 6-15. Amine sweetening verses membrane sweetening.

6.11.9 Membrane Advantages & Disadvantages
Recently, several very large membrane systems including some treating in excess of 300MMscfd of gas containing over 30% CO2 have been successfully implemented as a result of advances in pretreatment technology. Membrane system suppliers include Grace, Kvaerner Process Systems, Natco-Cynara,UOP, and Ube.
Natural gas sweetening and dehydration using membranes often offers significant advantages over the more conventional methods such as amine treating, physical solvents, and solid adsorbents.
Membranes are particularly attractive when the pressure of the feed gas is high (over 500 psig) and/or the CO2 content of the gas to be treated is high (over 10%).

6.11.9.1 Advantages
Lower Capital Cost Capital Expenditure (CAPEX)
Membrane systems are skid-mounted, except for larger pretreatment vessels.
Scope, cost, and time required for site preparation are minimal. Installation costs are significantly lower than alternative technologies, especially for remote areas and offshore installations.
Membrane systems do not require the additional facilities, such as solvent storage and water treatment needed by other processes.

Lower Operating Costs Operating Expense (OPEX)
The only major operating cost for single-stage membrane systems is replacement.
Cost is significantly lower than the solvent replacement and energy costs associated with traditional technologies. Improvements in membrane and pretreatment design allow a longer useful membrane life, which further reduces operating costs. Energy costs of multistage systems with large recycle compressors are usually comparable to those for traditional technologies.

Deferred Capital Investment
Gas flow rates often increase over time as more wells are brought online. With traditional technologies, the system design needs to take this later production into account in the initial design; thus, the majority of the equipment is installed before it is even needed. The modular nature of membrane systems means only the membranes that are needed at start-up need be installed. The rest can be added, either into existing tubes or in new skids, only when they are required. On offshore platforms, where all space requirements must be accounted for, space can be left for expansion skids rather than having to install them at the start of the project.

High Turndown
The modular nature of membrane systems means that low turndown ratios, to 10% of the design capacity or lower, can be achieved. Turn-up and turn-down increments can be set at whatever level is required during the design phase.

Operational Simplicity and High Reliability
Single-Stage Membrane Systems
Single-stage membrane systems have no moving parts. They have almost no unscheduled downtime. They are simple to operate. They can operate unattended for long periods, provided that external upsets, such as well shutdowns, do not occur. Equipment in pretreatment systems that could cause downtime, such as filter coalescers, are usually spared so that production can continue while the equipment is under maintenance. The addition of a recycle compressor adds some complexity to the system but much less than with a solvent- or adsorbent-based technology.
Multistage Membrane Systems
Multistage membrane systems can be operated at full capacity as single-stage systems when the recycle compressor is down, although hydrocarbon losses will increase. Start-up, normal operation, and shutdown of a complex multistage system can be automated so that all-important functions are initiated from a control room with minimal staffing.

Good Weight and Space Efficiency
Skid construction can be optimized to the space available. Multiple elements can be inserted into tubes to increase packing density. Space efficiency is especially important for offshore environments, where deck area is at a premium.

Adaptability
Because membrane area is dictated by the percentage of CO2 removal rather than absolute CO2 removal, small variations in feed CO2 content hardly changes the sales gas CO2 specification. For example, a system designed for 10% down to 3% CO2 removal produces a 3.5% product from a 12% feed gas and a 5% product from a 15% feed gas. Adjusting process parameters such as operating temperature, the designer can further reduce the sales gas CO2 content.

Environmental Friendly
Membrane systems do not involve the periodic removal and handling of spent solvents or adsorbents. Permeate gases can be flared, vented, or reinjected into the well or used as fuel. Items that do not need disposal, such as spent membrane elements, can be incinerated.

Design Efficiency
Membrane and pretreatment systems integrate a number of operations such as dehydration, CO2 and H2S removal, dew-point control, and mercury removal. Traditional CO2 removal technologies require all of these operations as separate processes and may also require additional dehydration because some technologies saturate the product stream with water.

Power Generation
Permeate gas from membrane systems can be used to provide fuel gas for power generation, either for a recycle compressor or for other equipment. This virtually free fuel production is especially useful in membrane-amine hybrid systems, where the membrane system provides all the energy needs of the amine system.

Ideal for De-bottlenecking
Because expanding solvent- or adsorbent-based CO2 removal plants without adding additional trains is difficult, an ideal solution is to use a membrane for bulk acid gas removal and leave the existing plant for final cleanup. An additional advantage is that the permeate gas from the membrane system can often be based as fuel for the existing plant, thus avoiding significant increase in hydrocarbon losses.

Ideal for Remote Locations
Many of the factors mentioned above make membrane systems a highly desirable technology for remote locations where spare parts are rare and labor unskilled. Solvent storage and trucking, water supply, power generation (unless a multistage system is installed), or extensive infrastructure are not required.

6.11.9.2 Disadvantages
Clean feed: Pretreatment of the feed to the membrane to remove particulates and liquids is generally required
Gas compression: Because pressure difference is the driving force for membrane separation, considerable recompression may be required for either or both the residue and permeate streams
Generally higher hydrocarbon losses than solvent systems.
Membrane materials are not suitable for high hydrogen sulfide partial pressures and applications for bulk H2S removal is not practical. According to Kvaerner Process Systems the maximum permissible H2S partial pressure is around 20 psia for present membrane materials.

6.11.10 Hybrid Configurations
In some situations, placing a single stage membrane system upstream of an amine unit has a very positive effect. The presence of one unit eliminates the shortcomings of the other and the
combined “hybrid” system becomes less expensive to build and operate and more flexible in handling changes in feed gas conditions. Here is a list of most of the potential benefits when using a hybrid system:


Operating Issues
Hybrid vs Amine Only Hybrid vs Membrane Only
Hydrocarbon Losses Increased losses, unless there is a use for the permeate Slight increase in losses, but typically no compression
Meets Low CO2 Spec Same Yes, much better
Meets Low H2S Spec Same Yes, much better
Energy Consumption Lower Higher
Operating Cost Lower Higher
Maintenance Cost Slightly higher Higher
Ease of Operation Slightly more complex More complex
Dehydration Product still saturated Re-saturates product gas
Corrosion Potential Lower (lower loadings) Not a concern
Amine Foaming Virtually eliminated Not a concern
Capital Cost Issues
Hybrid vs Amine Only Hybrid vs Membrane Only
Recycle Compression Not a concern Eliminates need for compression
Total Installed Cost Same to lower Higher
Very Large Gas Flow Significant savings Higher
Table 6-16. Comparison of Hybrid to Amine and Membrane CO2 Removal Systems
As these comparisons are very dependent upon the natural gas being processed, the operating conditions, the economic variables and the location of the processing facility. It is important to understand the areas where each technology in a hybrid system performs best.

The size of an amine unit is directly related to the number of moles of CO2 removed from the feed gas. As CO2 content rises from low to moderate partial pressures in the feed, the rich solvent loading increases to somewhat offset the increased demand for solvent. But as partial pressures increase to high levels, the solvent approaches a maximum loading. At that point, any increase in CO2 can only be removed by increasing the circulation rate. The same is not true for membranes. Permeation increases as the feed gas CO2 partial pressure increases, making the membrane much more efficient at high concentration of CO2.

As mentioned earlier (comparison table), meeting low CO2-content sales gas specifications causes single-stage membranes to lose efficiency, while amines work very economically. By combining the technologies in series to treat gases with a high partial pressure of CO2, the membrane operates where it is best (high concentrations of CO2) and the solvent system works where it is best (achieving low specification for treated gas CO2 content).
The obvious, and first, application of hybrid systems was in enhanced oil recovery (EOR). The CO2 content is extremely high, 70% or more, in these plants.
Clearly, high-CO2 natural gas streams are good candidates for using membranes to remove all or part of the acid gas.
Because membranes are more efficient for processing high partial pressures of CO2, the capital and operating costs are typically lower for a hybrid when compared to a solvent-only system. The issue that must be carefully monitored is the amount of hydrocarbon (specifically, methane) lost with the CO2 in the permeate stream.
Recently, a study was conducted for a plant with a design flow of 240 MMSCFD and an inlet CO2 content of 41%. The specification was 3% CO2 upstream of the cold plant to insure a 5% pipeline specification in the residue sales gas. Detailed cost estimates were developed for standalone amines and a hybrid unit. Stand-alone membranes were not considered due to customer preferences.

Examples of Working Hybrid Systems
There are many hybrid systems currently operating around the world. Data is not available on all of the applications, but some are presented here to demonstrate the ways in which hybrids are used.
Early applications of membrane technology were in the area of enhanced oil recovery with CO2 flood. In general, high CO2 content of a gas is a good indicator for the use of membranes and/or hybrid systems. As will be shown, low CO2 pipeline specifications are another reason to adopt a hybrid configuration.

Example (Plant) 1
The first plant is an example of how not to apply a hybrid system. The plant processed a very dry gas in west Texas. The initial cut of CO2 was made with a single stage membrane. The problem was the hydrocarbon loss in the membrane. Reducing CO2 from 55% to 7% in a single stage results in high losses. A better design would have been to remove less CO2 with the membrane and more with the amine.

Fig. 6-38 . Example 1 of hybrid acid gas treatment system.

Example (Plant) 2
The second plant has done a very good job of maximizing capacity. Half the gas is processed in a two-stage membrane to reduce CO2 from 21% to <5%. The remainder of the gas bypasses the membrane and is blended to obtain 13% CO2 going to the amine unit. Despite the deep cut taken by the membrane unit, hydrocarbon losses are reduced by using a two-stage membrane configuration. This unit has operated for more than 10 years with very little trouble.


Fig. 6-39 . Example 2 of hybrid acid gas treatment system.


Hybrid Economic Evaluation
To illustrate some advantages of hybrid systems, a recent evaluation is presented by UOP in which a large amount of gas is to be processed. In this case, the natural gas processor wants to remove CO2 upstream of an NGL recovery plant. Stand-alone solvent systems are compared to hybrid units, which pair a single-stage membrane with either hot potassium carbonate or a specialty amine using a low-energy flow scheme.
The designs for the stand-alone solvent systems reveal that they cannot be built economically in a single train configuration at the remote location. The limiting factor is the diameter of the absorber and regenerator. They exceed the width that can be accommodated during trucking equipment to the site. The hybrid configuration offers an opportunity to reduce the size of the solvent system while keeping all the equipment within the transportation limitations.
The equipment costs for each system were estimated and then an installed cost was determined by applying a multiple to the equipment cost. The installation factors were chosen based on typical costs for units of similar scope and size. Operating costs were estimated based on heat duty, solvent cost, membrane element replacement and the cost of electricity. The relative results are shown in Table 3.


CO2 Removal 20% to <2% STAND-ALONE HYBRID
Case 1 Case 2 Case 3 Case 4
Hot Pot Amine Hot Pot Amine
2 Trains 2 Trains 1 Train 1 Train
Capital Cost 1.1 1.2 1.0 1.1
Annual Operating Cost 1.9 1.7 1.0 1.0
Table 6.17. Cost Comparison for Stand-alone and Hybrid Systems

In this example, the hybrid case pairing a membrane unit and hot potassium carbonate was the least expensive option. The stand-alone hot potassium carbonate system was 10% higher in capital cost and 90% more expensive to operate. The table shows both single-train hybrid options to be lower in cost than either stand-alone two-train system.
It should be noted that this analysis did not assume any hydrocarbon loss for the hybrid systems. It was assumed that any methane in the permeate would be burned to produce reboiler heat in the solvent system and supply low-Btu gas into the site fuel header. If these options had not existed, operating cost would have increased significantly for the hybrid options
6.12 Biological Processes
A biological process for removing H2S from natural gas has been reported by Cline et al. (2003). In this process, the gas stream that contains the H2S is first absorbed into a mild alkaline solution, and the absorbed sulfide is oxidized to elemental sulfur by naturally occurring microorganisms.
Cline et al. (2003) reported the successful startup and operation of a commercial plant in Canada. The plant is designed to treat a high-pressure natural gas and produce a product that meets H2S specifications (4 ppmv or lower). The sulfur plant capacity is approximately 1 ton per day, with a sulfur removal efficiency of 99.5% or higher. Cline et al. (2003) offer the following claims regarding the process:
• Cost effective up to 50 tons/d (50 tonnes/d) of sulfur
• H2S concentrations in the sour gas from 50 ppmv to 100 vol%
• Sour gas pressures to 75 barg (1,100 psig)
• H2S concentrations in the sweet gas 1 ppmv
• Formation of hydrophilic sulfur that prevents equipment fouling or blocking

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Gas Sweetening and Sulfur Recovery - Chapter 6 - Part 3 B

6.13 Process Selection
6.13.1 Inlet Gas Stream Analysis
An accurate analysis cannot be overstressed. Process selection and economics depend on knowing all components present in the gas. Impurities, such as COS, CS2, and mercaptans (even in small concentrations), can have a significant impact on the design of gas sweetening processes and downstream processing facilities.
When sulfur recovery is required, the composition of the acid gas stream feeding the sulfur plant must be considered. When CO2 concentrations are >80%, selective treating should be considered to raise the H2S concentration to the sulfur recovery unit (SRU). It may involve a multistage treating system. High concentrations of H2O and hydrocarbons can cause design and operating problems in the SRU. The effect of these components must be weighed when selecting a gas sweetening process.
Process selection can often be based on gas composition and operating conditions. High acid gas partial pressures, 50 psi (345 kPa) and above, increase the likelihood of using a physical solvent. The presence of significant quantities of heavy hydrocarbons in the feed discourages the use of physical solvents.
Low partial pressures of acid gases and low outlet specifications generally require the use of amines for adequate treating.

6.13.2 General Considerations
Each treating processes has advantages relative to the others for certain applications.
When making a final selection, the following facts should be considered:
• Type of acid contaminants present in the gas stream
• Concentrations of each contaminant and degree of removal required
• Volume of gas to be treated and temperature and pressure at which the gas is available
• Feasibility of recovering sulfur
• Desirability of selectively removing one or more of the contaminants without removing the others

6.13.3 Removal of H2S
The presence and amount of heavy hydrocarbons and aromatics in the gas can affect the environmental conditions required at the plant site. Removal of H2S to meet pipeline qualities (4 ppm) can be categorized as follow:

Feeds with Small Acid Gas Loadings
Batch processes should be considered for feeds with small acid gas loadings. The most common processes include iron sponge, sulfa-treat and sulfacheck.

Feeds with Moderate to High Acid Gas Loadings
For feeds with moderate to high acid gas loadings, the disposal and replacement costs are high. There is a need to select a process that can be regenerated. Amine systems are most often used. DEA is the most commonly used amine. The acid gas stream coming off the amine stripper can be flared at moderate loadings or converted to elemental sulfur at higher loadings

Process Must Be Added Downstream of the Amine System
A process must be added downstream of the amine system to converts acid gas to sulfur. Processes commonly used include LO-CAT® and Claus. Some gas streams can be treated directly with LO-CAT® solution and thus, the need to separate the acid gas components from the gas stream with an amine unit is eliminated. When a Claus unit is used, it may be necessary to add tail gas cleanup downstream of the Claus unit if acid gas loadings are very high. This is normally accomplished with amine-based system because the acid gas from the stripper can be vented (assuming levels of H2S in the gas being treated are very low). Gas permeation is attractive for low volume gas streams in remote areas where the loss of methane is not critical.
Systems with a second-stage recycle may be competitive with amine systems.

6.13.4 Removal of H2S and CO2
Often both H2S and CO2 are present and must be removed. Essentially, all of the H2S will have to be removed and only a fraction of the CO2 will have to be removed.

Feeds with Low Concentrations of CO2
For feeds with low concentrations of CO2, it is usually economical to use a nonselective solvent such as MEA or DEA. These processes require equipment be sized to essentially remove all the CO2 so that the H2S specification can be achieved.

Feeds with Increasing Concentrations of CO2
For feeds with increasing concentrations of CO2, it is often economical to use a selective process such as MDEA, Sulfinol®, or Selexol®, which will remove a higher percentage of H2S than CO2 from a stream. Another alternative is to use gas permeation or a carbonate system for bulk removal of CO2 upstream of a nonselective amine unit.
It may be economical to remove both H2S and CO2 to a level where the CO2 content is acceptable with either a selective or nonselective process, and use a sulfur removal process (iron sponge, Sulfa-Treat, Sulfa-Check, LO-CAT®) for final treating of the residue gas.

6.13.5 Process Selection charts
Four scenarios are possible for acid gas removal from natural gas:
1. CO2 removal from a gas that contains no H2S
2. H2S removal from a gas that contains no CO2
3. Simultaneous removal of both CO2 and H2S
4. Selective removal of H2S from a gas that contains both CO2 and H2S

Because the concentrations of CO2 and H2S in the raw gas to be processed and the allowable acid gas levels in the final product vary substantially, no single process is markedly superior in all circumstances, and, consequently, many processes are presently in use.

Selection criteria for the solvent-based processes are discussed by Tennyson and Schaaf (1977) and Figure 6-40 to Figure 6-43 are based on their recommendations. The guidelines in these figures are naturally approximate and should be treated as such. These figures are for solvent-based processes only. Thus, they exclude some commonly used processes such as adsorption and membranes. Note that “hybrid” in the figures denotes mixed-solvent systems that contain both amine and a physical solvent.


Fig. 6-40. Process selection chart for CO2 removal with no H2S present.
http://oilprocessing.net/data/documents/V6-41.png
Fig. 6-41.Process selection chart for H2S removal with no CO2 present.

Fig. 6-42. Process selection chart for simultaneous H2S and CO2 removal.

Fig. 6-43. Process selection chart for selective H2S removal with CO2 present.

6.14 Safety & Environmental Considerations
6.14.1 Amines
The obvious safety concern with amine treating is the potential for H2S leaks in the plant, even from spilled rich amine. In addition, some sections operate at high temperatures. Caustic handling is another hazard if MEA reclaiming is performed on-site. Reclaimer waste products are toxic and must be handled with care.
From an environmental perspective, in addition to the remote chance of hydrogen sulfide release, amines have an affinity for BTEX (benzene, toluene, ethylbenzene, and xylenes), which may be vented during amine regeneration if the sulfur is not recovered. If MEA or DGA are used with reclaimers, the reclaimer solids present a disposal problem, especially with MEA because caustic or soda ash is added to help reverse the reactions.

6.14.2 Physical Absorption
When H2S or other sulfur compounds are removed from a gas stream that contains high levels of these materials, the potential always exists for a leak. Depending upon the process, the solvent may be hazardous or toxic.

6.14.3 Adsorption
Safety and environmental problems associated with adsorbents such as molecular sieve are relatively minor. Thorough regeneration and purging must be done before the adsorbent can be replaced. However, it should be nonhazardous and disposable in a land fill.
Most solid scavengers are respiratory and eye irritants. Spent iron sponge is pyrophoric, and great care must be taken in the removal and disposal of reacted iron-sponge material. The manufacturer’s recommendations for this material must be carefully followed to prevent a serious incident.

6.14.4 Membranes
Membranes are probably the safest and most environmentally friendly of the processes for gas treating. No chemicals are used, no waste disposal by-products are generated, and membranes operate at low pressures and generally ambient temperatures.

6.15 Design Procedure
6.15.1 Iron Sponge
Step 1
The minimum vessel diameter for gas velocity is given by:
dmin = 60 (QgTZ/PVg max)0.5 Eq. 6-41
where
dmin = minimum internal vessel diameter, in. Qg = gas flow rate, MMscfd;
T = operating temperature,0R. Z = gas compressibility factor
P = operating pressure, psia. Vg max =maximum gas velocity, ft/s.

Step 2
The maximum rate of deposition of 15 grains of H2S/min-ft2 of bed cross-sectional area is also recommended to allow for the dissipation of the heat of reaction (1 grain = 1.428 x 10-4 lb). The following establishes a minimum required diameter for deposition given by
dmin = 8945 (Qg x H2S /Ø ) 0.5 Eq. 6-42
where
ø = rate of deposition, grains/min ft3 . H2S = mole fraction of H2S.

Step 3
The larger of the diameter calculated by Equations (6-41) or (6-42) will set the minimum vessel diameter. At very low superficial gas velocities <2 ft/s, channeling of the gas through the bed may occur. An upper limit to the vessel diameter may be determined by the following equation assuming a minimum velocity of 2 ft/s:
dmax = 60 (QgTZ/PVg min)0.5 Eq. 6-43
where
dmax = maximum internal vessel diameter, (in.). Vg min = minimum gas velocity, ft/s.

Step 4
A contact time of 60 s is considered a minimum in choosing a bed volume.
A larger volume may be considered as it will extend the bed life and thus extend the cycle time between bed changes. Assuming a minimum contact time of 1 min, any combination of vessel diameter and bed height that satisfies the following is acceptable:
d2H ≥ 3600 (QgTZ / P) Eq. 6-44
where
d = vessel internal diameter, in. H = bed height, ft.

When selecting acceptable combinations, the bed height should be at least 10 ft (3 m) for H2S removal and 20 ft (6 m) for mercaptan removal.
The selection should produce sufficient pressure drop to ensure proper flow distribution over the entire cross-section. The vessel diameter should be between dmin and dmax.
The iron sponge is normally sold by the bushel.

Step 5
The amount of iron oxide, which is impregnated on the wood chips, is normally specified in units of pounds of iron oxide (Fe2O3) per bushel. Common grades are 6.5, 9, 15, or 20 lbs Fe2O3/bushel. Theoretical bed life for the iron sponge between replacements is determined from
tc = 3.14 x 10-8 Fe d2 H E/ (Qg x H2S) Eq. 6-45
where tc = cycle time, days; Fe = iron sponge content, lbs Fe2O3/bushel; E = efficiency (0.65-0..

Example 6-1
Iron Sponge Unit
Given
Qg = 2 MMscfd - SG = 0.6 - H2S = 19 ppm - P = 1200 psig (1214.7 psia)
T = 100 0F (560 0R) - Z = 0.85
maximum gas velocity = 10 ft/s - Use minimum gas velocity = 2 ft/s
Use a rate of deposition (Ø), of 15 grains/min-ft2 - Use cycle time, tc = 30 days;
Use Fe =iron sponge content, 9 lb Fe2O3/bushel; - Use E=efficiency (0.65)
Solution
Step 1. Calculate Minimum Vessel Diameter for Gas Velocity (Eq. 6-41)
dmin = 60 [(2 x 560 x 0.85) /(1214.7 x10)]0.5
dmin = 16.8 in.

Step 2. Calculate Minimum Vessel Diameter for Deposition (Eq. 6-42)
dmin = 8945 (2 x 0.000019 /15 ) 0.5
dmin = 14.2 in.

Step 3. Calculate Maximum Diameter (Eq. 6-43)
dmax = 60 (2 x 560 x 0.85/1214.7 x 2)0.5
dmax = 37.6 in.

Therefore, any diameter from 16.8 to 37.6 in. is acceptable.
Step 4. Choose a Cycle Time of 1 Month or Longer (Eq. 6-45)
30 = 3.14 x 10-8 x 9 x d2 H x 0.65/ (2 x 0.000019) Eq.4545
d2 H = 6206
Check for the minimum contact time (Eq. 6-44)
6206 ≥ 3600 (2 x 560 x 0.85 /1214.7)
6206 ≥ 2821 ok.
In a table, chose a diameter between 16.8 to 37.6 and calculate the corresponding height.

d (in.) H (ft)
18 19.1
24 10.8
30 6.9
36 4.8
Table 6-18. Solution of example 6-1
An acceptable choice is a 30 in. O.D. vessel. Since tc and E are arbitrary, a 10 ft bed is appropriate.
6.15.2 The Amine System
Method 1
Amine Circulation Rates
The circulation rates for amine systems can be determined from the acid gas flow rates by selecting a solution concentration and an acid gas loading. The following equations can be used:
LMEA = 112 Qg XA / c ρ AL Eq. 6-46

LDEA = 192 Qg XA / c ρ AL Eq. 6-47
where
LMEA = MEA solution circulation rate, gpm; - LDEA = DEA solution circulation rate, gpm;
Qg = gas flow rate, MMscfd); - XA = required reduction in total acid gas fraction, moles acid gas removed/mole inlet gas.
Note: XA represents moles of all acid components, that is, CO2, H2S, and meracaptans, as MEA and DEA are not selective;
c = amine weight fraction, kg amine/kg solution (lbs amine/lbs solution);
ρ = solution density, lbs/gal; - AL = acid gas loading, mole acid gas/mole amine.
The specific gravity of amine solutions at various amine concentrations is in Figure 6-44.
For design purposes, the following solution strengths and loadings are recommended to provide an effective system without excessive corrosion rates:
MEA solution strength—20 wt% MEA
DEA solution strength—35 wt% DEA
MEA acid gas loading—0.33 mol acid gas/mol MEA
DEA acid gas loading—0.5 mol acid gas/mol DEA
Density of MEA—8.41 lbs/gal
Density of DEA—8.71 lbs/gal
Using the recommended concentrations and specific gravities at 20 0C from Figure 6-44:
20%MEA = 1.008SG = 1.008 x 8.34 lbs/gal
= 8.41 lbs/gal = 8.41 x 0.20
= 1.68 lbs MEA/gal
=1.68/61.08 = 0.028 mol MEA/gal
35%DEA = 1.044SG = 1.044 x 8.34 lbs/gal
=8.71 lbs/gal = 8.71 lbs/gal x 0.35
= 3.05 lbs DEA/gal
= 3.05/105.14 = 0.029 mol DEA=gal
Using these design limits, the circulation rates required can be determined from Equations (6-48) and (6-49):
LMEA = 202 Qg XA Eq. 6-48

LDEA = 126 Qg XA Eq. 6-49
The circulation rate determined with the above equations should be increased by 10-15% to supply an excess of amine. The rates determined can be used to size and select all equipment and piping.

Amine Reboiler
The reboiler duty can be estimated as follows:
qerb = 72,000 LMEA Eq. 6-50
qerb = 60,000 LDEA Eq. 6-51
where
qreb = reboiler duty, (Btu/h) - LMEA = MEA circulation rate, gpm - LDEA = DEA circulation rate, gpm.
The operating temperature for amine reboilers is determined by the operating pressure and the lean solution concentration. Typical reboiler temperature ranges are as follows:
MEA reboiler— 225-260 0F (107-127 0C)
DEA reboiler—230-250 0F (110-121 0C)
For design purposes, the reboiler temperature for a stripper operating at 10 psig (69 kPa) can be assumed to be 245 0F (118 0C) for 20% MEA, and 250 0F (121 0C) for 35% DEA. Boiling point versus solution concentration curves at various pressures are shown in Figure 6-45.


Fig. 6-44. Specific gravity of amine solution versus composition.

Fig. 6-45. Boiling points of aqueous monoethanolamine & diethanolamine solutions at various pressures.

Example 6-2
Amine Processing Unit (DEA)
Given
Gas volume = 100 MMscfd - Gas gravity = 0.67 SG (air =1.0) - Pressure =1000 psig
Gas temperature = 100 0F - CO2 inlet = 4.03% - CO2 outlet = 2%
H2S inlet = 19 ppm = 0.0019% - H2S outlet = 4 ppm - Max. ambient temp.= 100 0F

Solution
Step 1. Process Selection
Total acid gas inlet = 4.03+0.0019 = 4.032%
Partial pressure of inlet acid gas = 1014.7 x (4.032/100) =40.9 psia

Total acid gas outlet = 2.0%
Partial pressure of outlet acid gas = 1014.7 x (2.0/100) = 20.3 psia

From Figure 6-40 (CO2 removal, no H2S present) for removing CO2 and H2S, possible processes are amines, and carbonates.

Step 2. DEA Circulation Rate
Determine the circulation rate (Equation 6-47):

ρ = DEA density, =8.71 lb/gal, c = 0.35 lb/lb; AL = 0.50 mole/mole; Qg =100 MMscfd; XA = 4.032% = 0.04032
Note: In order to meet the H2S outlet, virtually all the CO2 must be removed, as DEA is not selective for H2S.
LDEA = 192 x 100 x 0.04032 / (0.35 x 0.871 x 0.5) = 508 gpm
Add 10% for safety = 560 gpm.

Step 3. Reboiler Duty
Determine the reboiler duty (Equation 6-51):
qerb = 60,000 x 560 = 33.6 MMBtu/h

Method 2
A simplified procedure for making rough estimates of the principal parameters for conventional MEA, DEA and DGA® amine treating facilities when both H2S and CO2 are present in the gas is given below.
The procedure involves estimating the amine circulation rate and using it as the principal variable in estimating other parameters. For estimating amine circulation rate, the following equations are suggested:

For MEA:
GPM = 41 x (Qy/x) Eq. 6-52
(0.33 mol acid gas pick-up per mole MEA assumed)

For DEA (conventional):
GPM = 45 x (Qy/x) Eq. 6-53
(0.5 mol acid gas pick-up per mole DEA assumed)

For DEA (high loading):
GPM = 32 x (Qy/x) Eq. 6-54
(0.7 mol acid gas pick-up per mole DEA assumed)

For DGA®
GPM = 55.8 x (Qy/x) Eq. 6-55
(0.39 mol acid gas pick-up per mole DGA® assumed)
(DGA® concentrations are normally 50-60% by weight)

Where:
Q = Sour gas to be processed, MMscfd
y = Acid gas concentration in sour gas, mole%
x = Amine concentration in liquid solution, wt%

After the amine circulation has been estimated, heat and heat exchange requirements can be estimated from the information in Table. 6-19. Pump power requirements can be estimated from Table. 6-20.

Duty, Btu/hr Area, Sq ft.
Reboiler (Direct fired) 72,000 x GPM 11.30 x GPM
Rich-Lean Amine HEX 45,000 x GPM 11.25 x GPM
Amine cooler (air cooled) 15,000 x GPM 10.20 x GPM
Reflux condenser 30,000 x GPM 5.20 x GPM
Table 6-19. Estimated Heat Exchange Requirements

Main Amine Solution Pumps GPM x PSIG x 0.00065 = HP
Amine Booster Pumps GPM x 0.06 = HP
Reflux Pumps GPM x 0.06 = HP
Aerial Cooler GPM X 0.36 = HP
Table 6-20 . Estimated Power Requirements

Eqs 6-52 to 6-55 normally provide conservative (high) estimates of required circulation rate. They should not be used if the combined H2S plus CO2 concentration in the gas is above 5 mole%. They also are limited to a maximum amine concentration of about 30% by weight.
The diameter of an amine plant contactor, can be estimated using the following equation:

Dc = (1936 Q / p0.5 )0.5 Eq. 6-56

Where:
Q = MMscfd gas to contactor, - P = Contactor pressure is psia
Dc = Contactor diameter in inches before rounding up to nearest 6 inches.

Similarly, the diameter of the regenerator below the feed point can be estimated using the following equation:

Dr = 3.0 x (GPM)0.5 Eq. 6-57

Where:
GPM = Amine circulation rate in gallons per minute, - Dr = Regenerator bottom diameter in inches
The diameter of the section of the still above the feed point (Dra) can be estimated at 0.67 times the bottom diameter.
Example 6-3
30.0 MMscfd of gas available at 850 psig and containing 0.6% H2S and 2.8% CO2 is to be sweetened using 20%, by weight, DEA solution. If a conventional DEA system is to be used, what amine circulation rate is required, and what will be the principal parameters for the DEA treating system?

Solution:
Using Eq 21-7, the required solution circulation is:
GPM = 45(Qy/x) = 45(30 x 3.4/20) = 230 gallons of 20% DEA solution per minute.

Heat exchange requirements (from table 6-20)
Reboiler
H = 72,000 x 230 = 16.6 x 106 Btu/hr
A = 11.3 x 230 = 2600 ft2
Rich-Lean amine exchanger
H = 45,000 x 230 = 10.4 x 106 Btu/hr
A = 11.25 x 230 = 2590 ft2
Amine cooler
H = 15,000 x 230 = 3.45 x 106 Btu/hr
A = 10.2 x 230 = 2350 ft2
Reflux condenser
H = 30,000 x 230 = 6.9 x 106 Btu/hr
A = 5.2 x 230 = 1200 ft2

Power requirements (table 6-21)
Main amine pumps
HP = 230 x 850 x 0.00065 = 127
Amine booster pumps
HP = 230 x 0.06 = 14
Reflux pumps
HP = 230 x 0.06 = 14
Aerial cooler
HP = 230 x 0.36 = 83
Contactor diameter
Dc = (1936 Q / p0.5 )0.5
Dc = 44.5 inches or 48 inches rounded up.

Regenerator diameter below feed point:
Dr = 3.0 x (230)0.5
Dr = 45.5 inches or 48 inches (bottom) rounded up to nearest 6 inches.
Regenerator diameter above feed point:
Dra = 0.67 x 48 = 32.2 inches or 36 inches (top) rounded up to nearest 6 inches.

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Hydrocarbon Recovery - Chapter 7

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 7 297
Hydrocarbon Recovery 297
7.1 Introduction 297
7.2 Process Components 297
7.2.1 External Refrigeration 297
7.2.2 Turboexpansion 300
7.2.3 Heat Exchange 302
7.2.4 Fractionation 303
7.3 Hydrocarbon Recovery Processes 304
7.3.1 Dew Point Control and Fuel Conditioning 305
7.3.3 High Ethane Recovery 310

-------------
Chapter 7

Hydrocarbon Recovery

7.1 Introduction
To recover and separate NGL from a bulk of gas stream, a change in phase has to take place. In other words, a new phase has to be developed for separation to occur. (Chapter 2).

Pipeline quality natural gas specifications include limits on sulfur and water content, along with higher heating value, which must be about 950 to 1,150 Btu/scf (35,400 to 42,800 kJ/Sm3). Exact limits are set by negotiation between the processor and the purchaser. The previous chapters
addressed water, CO2, and sulfur specifications. This chapter addresses the heating value.
Unless the treated gas contains high concentrations of inert gases (N2, CO2), the heating value may be too high because of the C2+fraction present. This chapter discusses hydrocarbon recovery methods to both lower the heating value and create, simultaneously, valuable NGL liquid hydrocarbon products (Chapter 1).
Processors have additional reasons for reducing the C2+ fraction. Hydrocarbon recovery frequently is required in field operations for fuel conditioning or dew point control. As noted earlier, raw gas usually is usually too rich, and simple systems are usually used to lower the heating value (i.e., condition the fuel) by removing heavier hydrocarbons.
Dew point control (or “dew pointing”) is necessary when raw gas lines are constrained in liquid content as the liquid reduces gas throughput, causes slugging, and interferes with gas metering.
Dew point control is also necessary if a potential for condensation is present in a process because of temperature or pressure drops. The latter happens when the gas is in the retrograde condensation region. However, effective dew point control is much less demanding than C2+ recovery, as it can be accomplished without removal of a large portion of the C3+ fraction.

7.2 Process Components
The process elements involved in hydrocarbon recovery vary, depending upon the desired products and gas volume being processed as well as inlet composition and pressure.

7.2.1 External Refrigeration
External refrigeration is used to cool the gas stream to recover a significant amount of C3+ and to lower gas temperatures as the gas goes into other stages of hydrocarbon recovery. It may be the only source of refrigeration when inlet pressure is low. Adsorption and vapor compression refrigeration are used in special situations. However, vapor compression using propane as the working fluid is the most common in gas plants. (LNG facilities also use ethane or ethylene as well as hydrocarbon mixtures as refrigerants.

7.2.1.1 Propane Refrigeration Process
The refrigeration cycle consists of four steps that are depicted on the pressure− enthalpy chart in Figure 7.1:
1. Compression of saturated refrigerant vapor at point A to a pressure well above its vapor pressure at ambient temperature at point B
2. Condensation to point C by heat exchange with a cooling fluid, usually air.
3. Expansion through a valve (Joule-Thomson expansion) to cool and condense the refrigerant to point D
4. Heat exchange with the fluid to be cooled by evaporation of the refrigerant back to point A.

Figure 7.2 shows the flow diagram for a single-stage propane refrigeration system, with typical operating conditions. Each of the steps is described below.


Fig. 7-1. Schematic of refrigeration cycle on a pressure−enthalpy chart.

Fig. 7-2. Single-stage propane refrigeration system.

Compression Step: Cycle analysis begins with propane vapor entering the compressor as a vapor at 14.5 psia (1 bar) and approximately −400F (−400C), where it is compressed to 250 psia (17 bar) (point A to point B in fig. 7.1.
Condensation Step: The warm gas goes to an air- or water-cooled condenser, where the propane cools to 100 to 1200F (38 to 500C), totally condenses, and collects in a receiver (point B′ to point C in Figure 7.1).
Expansion Step: Propane liquid leaves the receiver and flashes through a J-T valve, where the temperature and pressure drop to −400F (−400C) and 16 psia (1 bar) (point C to point D). No change occurs in the enthalpy, but the temperature drops to the saturation temperature of the liquid at the expansion-discharge pressure.
Refrigeration Step: The cold propane then goes to a heat exchanger, the chiller, where it cools the process stream by evaporation (point D to point A in Figure 7.1).
Because the propane in the chiller is evaporating, and a minimal heat exchange occurs between cold propane vapor and the inlet gas, the inlet and outlet propane temperature remains constant. The propane returns to the compressor suction slightly above −400F (−400C).
In the past, most chillers were the kettle type in which propane is on the shell side and the liquid level is maintained above the tube bundle. Now, other high performance heat exchangers (e.g., plate-fin) are used. The chiller typically has two zones of heat transfer. The first is exchange of boiling propane with gas above its dew point and will involve only sensible heat. The second zone has condensing vapors from the process stream and boiling propane, which gives a much higher overall heat-transfer coefficient. To complete the cycle, the propane vapors leave the chiller and go to the suction drum before being compressed again.


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7.2.1.2 Alternate Process Configurations
Thermodynamics dictate that to minimize the refrigeration work (i.e., compression required) heat from the chiller should be removed at as high a temperature as possible. One way to reduce compressor duty per unit of refrigeration duty is to multistage the refrigeration process by removal of process heat at more than one stage. Figure 7.3 shows a two-stage system, with representative operating conditions. In this system, the condensed propane stream expands to about 62 psia (4.3 bar) and 25°F (−4°C) in two parallel J-T valves. One of the expanded streams goes to a chiller before going to the suction surge drum for the second stage of the compressor.
Vapor from the surge drum goes to the second stage of the compressor, while the liquid goes through a second J-T expansion to provide the low-temperature cooling. Pressures are adjusted so that the compression ratio is equal in both compressor stages. Table 7.1 shows the significant savings in going to two stages.
Multistaging reduces work requirements by removing heat from the process stream at different temperatures. An alternative is removal of heat from the refrigerant before it is expanded. Refrigerant sub-cooling is sometimes used by exchange of the propane that leaves the receiver with a portion of the cold liquid propane.
About 10% reduction in recirculation rate to the second stage of two-stage refrigeration is obtainable when the propane from the condenser is cooled 10°F (5°C).

Number of Stages 1 2 3
Change in compressor power (%) 0 -19.2 -23.3
Change in condenser duty (%) 0 -8.2 -9.6

Table. 7.1. Effect of Multistaging on Condenser and Compression Duty for Constant Refrigeration Duty with Propane as the Refrigerant. Refrigeration temperature is −40°F and condensing temperature is 100°F.

Although a second heat exchanger provides the most benefit from staging a refrigeration system, savings can be obtained by use of only a single J-T expansion into the economizer or by use of single J-T expansion and the heat exchanger. Gas plants commonly omit the first-stage heat exchanger for cost purposes. However, when energy requirements are critical, such as in an LNG plant, the J-T valve that bypasses the heat exchanger is commonly eliminated and three stages of refrigeration are common.
Another way to alter the temperature at which heat is removed is to use refrigerant cascading. In this option, one refrigerant (e.g., propane) is used to remove heat from another refrigerant (e.g., ethane), which then cools the process gas down to as low as −120°F (−85°C). This technique is commonly used in LNG plants. Many gas plant process stream temperatures go well below −120°F (−85°C) but usually rely upon process gas expansion for cooling instead of additional external cooling.

Fig. 7.3 Two-stage propane refrigeration system, with second heat exchanger and economizer. Units may omit either the first stage heat exchanger or expansion directly to the economizer.

Obviously many combinations are possible for making refrigeration more efficient. However, each must be balanced with the associated additional capital cost, realizable operating cost savings, and operating complexity. Although little incentive exists for more complex systems in gas processing, LNG plants make extensive use of complex refrigeration cycles to reduce refrigeration (i.e., compression) costs.

7.2.2 Turboexpansion
The J-T valve, which is essentially a control valve with a variable or fixed orifice, is an extremely simple, inexpensive, and widely used means to reduce gas temperature. Although still extensively used in many applications to produce refrigeration, J-Ts are being widely supplanted by turboexpanders in gas plants for cooling the process stream when it is a gas. Turboexpanders are, in essence, centrifugal compressors that run backwards. Unlike J-T expanders, they perform work during the process. Whereas J-T expansion is essentially an isenthalpic process (i.e no change in enthalby across the J-T valve, therefore, no work is done on or by the gas), an ideal, thermodynamically reversible turboexpander is isentropic (no heat transfer). The maximum reversible work required for compression is isentropic, and, conversely, the maximum reversible work recovered by a turboexpander system on expansion is also isentropic. Turboexpansion provides the maximum amount of heat removal from a system for a given pressure drop while generating useful work. The work is used to drive compressors or electrical generators.
The major breakthrough for turboexpanders came when the design and materials made it possible for condensation to occur inside the expander. The fraction condensed can be up to 50% by weight. However, the droplets must generally be 20 microns in diameter, or less, as larger droplets cause rapid erosion of internal components.
Most turboexpanders drive centrifugal compressors to provide a portion of the outlet compression. In situations where inlet pressures are very high (e.g., offshore) turboexpanders are used in pressure letdown to provide refrigeration for dew point control and to generate power.
Like compressors, expanders can be positive displacement or dynamic; dynamic can be radial or axial. Reciprocating expanders were used for liquefying gases. However, the only type used in gas processing is the radial unit with inward (centripetal) flow, and discussion is restricted to this type. A cutaway view of a typical turboexpander for gas processing is shown in Figure 7.4. The expander is the unit on the right, and the compressor is the unit on the left.
Gas enters the expander through the pipe at the top right, and is guided onto the wheel by the aerodynamically shaped adjustable guide vanes, which surround the expander wheel. The swirling high-velocity inlet gas turns the wheel and transfers part of its kinetic energy to the wheel and shaft, and exits to the right through the tapered nozzle. Because part of the energy of the gas has been transferred to the wheel, the exit gas is at a much lower temperature and pressure than the gas entering. The expander wheel, directly coupled to the compressor wheel, provides the work necessary to drive the centrifugal compressor on the left. Low-pressure gas enters in the straight section on the left, is compressed by the compressor wheel, and exits at the top of the unit. Lubricating oil enters in the top port shown in the center of the unit.


Fig. 7-4 Cutaway view of a turboexpander.

Operating conditions for turboexpanders vary, depending on the process application and the composition of the gas being processed. However, Table 7.2 provides a general idea of turboexpander operations in a gas plant.
Expander Compressor
Inlet gas rate, MMscfd (MMSm3/d) 250 (7.1) 187 (5.3)
Molar mass 22.80 19.69
Inlet pressure, psia (bar) 1,590 (109.7) 855 (58.9)
Inlet temperature, °F (°C) 30 (−1.0) 34 (1.0)
Outlet pressure, psia (bar) 870 (60.1) 1,165 (80.3)
Outlet temperature, °F (°C) −15 (−26.1) 78 (25.4)
Liquid formation, wt% 36 -
Power, kW 2,004 1,986
As tested efficiency, % 85.0 78.5
Speed, rpm 16,700 16,700
Wheel diameter, in. (mm) 7.75 (197) 10.9 (277)

Table 7.2. Expander−Compressor Design-Point Conditions
In principle, both the turboexpander and compressor can be multistaged. However, to date, mechanical sealing problems have made multistaging impractical for both the turbine and expander. The Engineering Data Book emphasizes some points that should be kept in mind for turboexpanders:
• Entrainment. Gas that enters the turboexpander must be free of both solids and liquids. Fine-mesh screens are used to protect the device, and the pressure drop across the screen should be monitored.
• Seal gas. This gas isolates process gas from the lubricating oil, or isolates process gas from the shaft if magnetic bearings are used, and must be clean and constantly available at the operating pressure. Sales gas is commonly used. Otherwise, a warmed inlet gas stream off of the expander inlet separator is used. (The gas must be warmed to 70°F [20°C] or more to prevent thickening of the lube oil, if used.)
• Lubricant pumps. These pumps must maintain a constant flow to lubricate the bearings if oil is used. A spare pump is mandatory.
• Shut-off valves. A quick-closure shut-off valve is used to shut in the inlet for startup and shutdown.
As is the case for centrifugal compressors, turboexpander efficiency diminishes when operating off of the design point. This variance can be about 5 to 7 percentage when the flow increases or decreases by 50%. However, the turboexpander normally is driving a compressor, which also will suffer loss in efficiency when off of the design point. Therefore, the overall effect on the turboexpander−compressor unit efficiency will be larger.

7.2.3 Heat Exchange
Most heat exchangers in a gas plant operating at or above ambient temperature are conventional shell and tube type and are ideal for steam and hot oil systems where fouling occurs. They are relatively inexpensive and easy to maintain because the tube bundle can be removed and tubes cleaned or replaced as needed.
Where the fluids are clean and fouling does not occur, such as in gas−gas exchangers, compact heat exchangers are ideal. This section briefly discusses two kinds, brazed-aluminum plate-fin heat exchangers and printed circuit heat exchangers, which commonly are used in gas processing.

7.2.3.1 Plate-Fin Exchangers
Cryogenic facilities have made extensive use of brazed-aluminum plate-fin heat exchangers since the 1950s. Instead of a shell and tube configuration, these units consist of channels formed by a thin sheet of aluminum pressed into a corrugated pattern (the fin) sandwiched between two aluminum plates. Each layer resembles the end view of corrugated cardboard. The fin channels may be straight or may have a ruffled or louvered pattern to interrupt the straight flow path.
Advantages of plate-fin exchangers include:
• Light weight.
• Excellent mechanical strength at subambient temperatures (used in liquid helium service [−4520F (−2680C)]). Can operate at pressures up to 1,400 psig (96 barg).
• High heat transfer surface area. Up to six times the surface area per unit volume of shell and tube exchanger and 25 times the area per unit mass.
• Complex flow configurations. Can handle more than 10 inlet streams with countercurrent, crossflow, and counter crossflow configurations.
• Close temperature approaches. Temperatures of 3°F (1.7°C) for single-phase fluids compared with 10 to 15؛F (6 to 9؛C) for shell and tube exchangers and 5°F (2.8°C) for two-phase systems.
Drawbacks and limitations of the exchangers include:
• Single-unit construction. Repair can be more costly and time consuming than with shell and tube exchangers.
• Maximum operating temperature of approximately 150°F (~85°C), although special designs go to 400°F (205°C).
• Narrow channels. More susceptible to plugging, and fine mesh screens are needed where solids may enter. Components that might freeze out, water, CO2, benzene, and p-xylene, must be in sufficiently low concentrations to avoid plugging. The exchangers can be difficult to clean if plugging occurs.
• Less rugged. Does not accept rough handling or high pipe stress on nozzles.
• Limited to fluids noncorrosive to aluminum. Caustic chemicals are corrosive but not corroded by acid gases, unless free water is present.
• Susceptible to mercury contamination. Mercury amalgamates with aluminum to destroy mechanical strength.
• Susceptible to thermal shock. Maximum rate of temperature change is 4°F/min (2°C/min), and maximum difference between two streams is 55°F (30°C).

7.2.3.2 Printed Circuit Heat Exchangers
Another heat exchanger type, the printed circuit heat exchanger (PCHE) is used in clean service. This technology is relatively new, commercialized in the 1980s, but hundreds of units are in service. Like electronic printed circuits, heat transfer passages are etched in plates, and the plates are bonded together by diffusion bonding. Unlike the brazed-aluminum exchangers, they are rugged and, depending on materials of construction, go to high temperatures and pressures but can still handle complex flow schemes that involve many streams. Heat transfer passage sizes range from “microchannels” (less than 8 mil, 200 microns) to “minichannels” (0.12 in, 3 mm) to provide high heat transfer surface areas. Heat transfer area per unit volume can be 800 compared with 500 for plate-fin exchangers. Offshore operations employ PCHEs in many applications because they offer comparable heat transfer at comparable pressure drops at significantly less size and at one fifth the weight.

7.2.4 Fractionation
In addition to conventional distillation columns, two other types of distillation columns are commonly found in gas plants: stabilizers and demethanizers. Stabilizers are stripping columns used to remove light ends from NGL streams. Demethanizers are also stripping columns to remove methane from the NGL bottoms product. Demethanizers also act as the final cold separator, a collector of cold NGL liquids, and source of recovering some refrigeration by cooling warm inlet streams.

7.2.4.1 Stabilizers
The primary focus of dew pointing or fuel conditioning is to obtain a leaner gas. However, the “by-product” is a liquid phase that contains a substantial amount of volatiles. To make the liquid product easier to store and to recover more light ends for fuel or sales gas, many of the systems will “stabilize” the liquid by passing it through a stabilizer column. The stabilizer feed typically enters at the top of a packed or tray column and no reflux occurs. To increase stripping of light ends, the column pressure will be lower than that of the gas separator that feeds the column. In some cases, a stripping gas may be added near the bottom of the column in addition to the externally heated reboiler installed to provide additional vapor flow and enhance light-ends removal. This feature usually comes as an increased operating cost because the gas from the stripper is at low pressure and must be recompressed if put back into the inlet gas stream upstream of the gas treating unit.

7.2.4.2 Demethanizer
A distinguishing feature of gas plants with high ethane-recovery rates is the demethanizer. The column differs from usual distillation columns in the following ways:
• It has an increased diameter at the top to accommodate the predominately vapor feed to the top tray.
• It is typically primarily a stripping column, with no traditional condenser–reflux stream.
• It may have several liquid feed inlets further down the column that come from low-temperature separators.
• It has a large temperature gradient; over 170°F (75°C) is common.
The column serves two main functions: it acts as a flash drum for the top feed, which comes in as a cold, two-phase stream, and it removes methane from the bottoms product. Depending upon the plant configuration, the feed may be from a turboexpander, a J-T valve, or a heat exchanger. In some configurations the columns have reflux, but many demethanizers have no reflux.
The NGL bottoms product is usually continuously monitored for methane content, which typically is kept below 0.5 liquid vol% of the ethane, on a C3+ free basis.
The top of the column usually operates in the −175 to −165°F (−115 to −110°C) range, with pressures in the 200 to 400 psig (14 to 28 barg) range.

7.3 Hydrocarbon Recovery Processes
Many process configurations are used to recover hydrocarbons in the field and in gas plants. The best configuration depends upon many variables, including:
• Product slate
• Gas volumes
• Gas composition
• Pressures, both inlet and outlet
The product slate dictates the required lowest operating temperature of the gas. Both dew point control and fuel conditioning have the same main product (a residue gas with reduced C3+ fraction). Dew point control is usually a field operation, and stabilization of the produced liquid is site specific. Although gas temperatures in a low-temperature separator (LTS) may go down to −40°F (−40°C), only a cold separator is required to separate the light ends from the liquid. Two new technologies, Twister and vortex tube, discussed below make dew point control and fuel conditioning a one-step process. In addition, use of membranes for fuel conditioning is discussed briefly.
If limited ethane recovery (<60% ethane) is desired, the recovery process is essentially a low temperature separator, except that fractionation of the cold liquid is added to increase the recovery. Lean oil absorption is sometimes used for up to approximately 50% ethane recovery. For high ethane recovery, the gas processing temperatures must be as low as −160°F (−110°C) and usually require a combination of external refrigeration and expansion. These plants require a demethanizer to increase recovery rates and to strip methane from the NGL.
Gas volumes and gas composition set the optimal plant configuration on an economic basis. This combination makes it difficult to set criteria for establishing the best plant configuration. However, the higher the gas volume and GPM (gal liquid per Mscf), the more attractive are high ethane recoveries.
Inlet gas pressures make a major difference in plant configuration. High pressures permit use of expansion J-T or turboexpander, to provide all of the cooling if low ethane recovery is desired. For low inlet pressures, either external refrigeration or inlet compression followed by expansion is needed to cool the gas, regardless of extent of ethane recovery. Required outlet pressure helps decide which approach should be taken.
The following three sections discuss the simple configurations of the three hydrocarbon-recovery systems:
1. Dew point control and fuel conditioning
2. Low ethane recovery
3. High ethane recovery

7.3.1 Dew Point Control and Fuel Conditioning
Dew point control and fuel conditioning exist to knock out heavy hydrocarbons from the gas stream. These operations are primarily field operations.

7.3.1.1 Low Temperature Separators
Low temperature separators (LTS) (also called low temperature extraction units, or LTX) are used both onshore and offshore. The process consists of cooling and partial condensation of the gas stream, followed by a low temperature separator.
When inlet pressures are high enough to meet discharge-pressure requirements to make pressure drop acceptable, cooling is obtained by expansion through a J-T valve or turboexpander. Otherwise, external cooling is required. Water usually is present, and to prevent hydrate formation the separator downstream of the expander is warmed above the hydrate-formation temperature to prevent plugging. An alternative to heating is injection of either ethylene glycol or methanol, which is then recovered and dried for reuse.
Figure 7.5 shows a LTS that uses ethylene glycol injection for hydrate prevention and uses J-T expansion for cooling. The feed initially goes through a water knockout vessel to remove free water. The water-saturated hydrocarbon gas and liquids then mix with ethylene glycol before being precooled. The mixture then passes through a J-T expansion valve and flashes into the low temperature separator to separate the gas, condensate, and glycol−water phases. The condensate goes to the condensate stabilizer for removal of remaining light ends.
Overhead gas from the low temperature separator passes through the precooler before being combined with the stabilizer overhead and put into the pipeline. The low temperature separator is set to maintain the proper dew point of the blended outlet gas.
The C3+ condensate from the stabilizer goes to product storage. The glycol−water mixture from the low temperature separator goes to the glycol regenerator for removal of the water and then reinjection into the feed.
If inlet pressures are too low for expansion, the stream is cooled by propane refrigeration. The advantage of direct refrigeration is that the pressure drop is kept at a minimum. Hydrate formation must be considered with either feed dehydration upstream of the unit or inhibitor injection. Glycol injection is usually the more cost effective, but if used, it increases the required refrigeration duty.

7.3.1.2 Twister
Twister, is a new device (from 1997) used for dew point control and dehydration. Figure 7.6 shows a cutaway view of the device and denotes the salient parts of the unit. Gas enters and expands through a nozzle at sonic velocity, which drops both the temperature and pressure and causes droplet nucleation. The two-phase mixture then contacts a wing that creates a swirl and forces separation of the phases by centrifugal force. The gas and liquid are separated in the diffuser; the liquid is collected at the walls and dry gas exits in the center.
Advantages of the system include:
• Simplicity. No moving parts and no utilities required.
• Small size and low weight. A 1-inch (24-mm) throat diameter, 6 feet (2 m) long tube can process 35 MMscfd (1 MMSm3/d) at 1,450 psia (100 bar).
• Driven by pressure ratio, not absolute pressure.
• Relatively low overall pressure drop. System recovers 65 to 80% of original pressure.
• High isentropic efficiency. Efficiency is around 90% compared with 75 to 85% for turboexpanders.
Drawbacks of the system include:
• Requires a clean feed. Solids erode the tubing and wing, necessitating an inlet filter separator.
• Limited turndown capacity. Flow variability is limited to +/- 10% of designed flow. This limitation is mitigated by use of multiple tubes in parallel.
• Liquids exit with a “slip gas.” This mixture is typically 20 to 30% of the total flow volume. The mixture can go to a gas−liquid separator for recovery of the gas, which may require recompression.


Fig. 7.5 Low-temperature separator (LTS), with glycol injection and condensate stabilization.


Fig. 7.6 Cutaway view of Twister device. (Courtesy of Twister BV.)

7.3.1.3 Vortex Tube
Vortex tubes use pressure drop to cool the gas phase but generate both a cold and warm gas stream. If streams are recombined, the overall effect is comparable to a J-T expansion. The principle of operation is the Ranque-Hilsch tube, developed in the 1940s and commonly marketed as a means to provide cold air from a compressed air stream.
For dew point control, and dehydration, the device has the vortex tube and a liquid receiver connected to the tube. Gas enters the tube tangentially through several nozzles at one end of the tube, expands, and travels spirally at near sonic velocities to the other end. As it travels down the tube, warm and cool gas separate. The cool gas goes into the center of the tube. Warm gas vents in a radial direction at the end, but the cool gas is reflected back up the tube and exits just beyond the inlet nozzles.
Condensation occurs in the cool gas, and the liquid is moved to the walls by centrifugal force, where it collects and drains into the receiver below. The overall cooling effect is comparable to that of a J-T expansion, with a low-temperature separator. However, the vortex tube combines the expansion and separation into a single step. The working pressure of the tube is 500 to 3,050 psig. The turndown ratio is 15% for a single tube but can be increased by use of multiple tubes in parallel; the optimum pressure drop is 25 to 35%. The vendor states that liquid condensation must be less than 10 wt%. The device performs well with up to 5% liquids in the inlet stream.
The device has been used to dehydrate gas from underground storage. To prevent hydrate formation in the cold stream, TEG is added. Like Twister, the vortex tube has the advantage of simplicity and light weight. It could be useful where limited turndown is acceptable. It will be of most value when no compression is required.

7.3.1.4 Membranes
As discussed in Chapter 6, membranes are being used in several areas of gas processing, including dew pointing. Membranes are ideal for this application, provided preconditioning is adequate to protect the membrane, and little penalty exists for permeate compression. Figure 7.7 shows the flow configuration. Gas enters the membrane on the discharge side of the compressor, and the residual gas provides fuel to the compressor engine or turbine. The low pressure permeate is recycled to the suction for recompression to recover the permeate. Table 7.3 provides results for one field unit. Gas rates are low because only a slip stream needs to be processed for fuel.
Like the previous two technologies, the process is simple and requires no moving parts. It too has the advantage of being relatively small and light weight. The technology is used on several offshore installations. Unlike the Twister and vortex tube, membranes have the advantage of a turndown ratio down to 50%, with no performance penalty. This property may not be an advantage for fuel gas conditioning, where flow rates should be stable.
The table points out the selectivity of the membrane and may be poorer than that of the above two technologies. In fuel conditioning, the selectivity is not a major issue because of the relatively small fraction of gas that needs to be recompressed, and the enriched stream is recycled without requiring additional compression.
Chapter 6 points out that membrane permeability is the product of the solubility and diffusion coefficient. For separation of light gases, the primary mode of selectivity is the diffusion coefficient. These membranes are silicone rubber compounds that preferentially absorb the heavy components.

7.3.2 Low Ethane Recovery
The focus of the previous section was removal of heavy components (C3+) to avoid condensation or to lower the heating value. This section discusses processes used in conventional gas plants, where the objective is to produce a lean gas and recover up to approximately 60% of the ethane in the feed gas. Two process schemes are used to obtain this level of ethane recovery:
• Cooling by expansion or external refrigeration
• Lean-oil absorption
As noted above, inlet pressure dictates the best means of refrigeration. Lean oil was an early method used for hydrocarbon recovery but is now used on a more limited basis. Many of the refrigerated lean oil absorption plants in operation today are large facilities, where replacing them with a more modern turboexpander plant would be capital cost prohibitive. Both approaches are described briefly below.


Fig. 7-7. Schematic for membrane unit used as a fuel conditioner.

Membrane Feed Conditioned Fuel Gas Permeate
Temperature, °F (°C) 95 (35) 51 (10.5) —
Pressure psig (barg) 940 (65) 940 (65) —
Total mass flow lb mol/h
(kg-mol/h) 110.1 (50) 58.0 (26.3) 52.1 (23.7)

Total volume flow MMscfd (MSm3/d) 0.95 (27) 0.5 (14) —
Mol% % Removed
Component Feed Fuel Gas Permeate from Gas
Carbon dioxide 1.3 0.6 2.08 76
Methane 72.8 81.2 63.59 41
Ethane 9.6 9.0 10.29 51
Propane 9.9 7.1 13.04 62
i-Butane 2.4 0.8 4.19 82
n-Butane 2.5 0.9 4.29 81
n-Pentane 1.3 0.4 2.30 84
Water 0.11 0.00 0.23 100
Hydrocarbon dew point (0C) 35 3.5 —


Table 7-3. Operating Conditions & Composition of NG Stream Using Membrane for Fuel Gas Conditioning

7.3.2.1 Cooling by Expansion or External Refrigeration
A general rule is to assume that ethane recovery increases with increased richness of the gas. This assumption is made because the ethane content in the vapor at the top of the column is set by column feed composition, along with temperature and pressure.
At constant pressure and temperature, the ethane concentration in the liquid decreases with increasing C3+ fraction, which lowers the ethane concentration in the vapor and, thus, increases the percent ethane recovered. (However, this outcome will not always be the case in plants that use J-T or turboexpanders, because leaner gas puts less of a load on the refrigeration-expander system and may lower column temperatures and increase recovery).

Figure 7-8 shows one commonly used direct-refrigeration process that employs recycle from a fractionator to maximize liquids recovery. Inlet gas is initially cooled with cold residue gas and cold liquid from the cold separator before going to the propane chiller and to the cold separator. Vapor from the separator is the sales gas, and the liquid goes to a fractionator to strip out light ends and recover liquid product.
The column operates at a lower pressure than does the cold separator. Because of system pressure drop and because the fractionator runs at the lower pressure, the recycle stream must be recompressed. Alternatives to the process include:
• Reduction or elimination of the recycle by adding reflux to the fractionator
• Running the fractionator at a higher pressure and use of a pump to feed the column from the cold separator

Fig. 7-8. Schematic of a direct refrigeration process for partial recovery of C2+ fraction.

These configurations assume that the gas enters sufficiently dehydrated to prevent hydrate formation. If the water content is higher, ethylene glycol can be added, which increases refrigeration duty. However, temperatures then are limited by glycol viscosity.
Because the unit relies only on external propane refrigeration, the lower temperature limit on the feed to the cold separator is −35°F (−37°C) at best. Unless the feed has a very high GPM, ethane recoveries will be below 60%. Expansion is required to lower temperatures and increase recoveries. With high inlet gas pressures, replacing the propane system with an expander is an attractive option. However, inlet compression may be necessary to obtain the temperatures required to obtain the desired recoveries. Both J-T and turboexpandersare used.
Crum (1981) points out that the J-T system may be preferable to turboexpanders, although recent advances in turboexpander technology may temper some of them:
• Low gas rates. J-T is more economically viable at low gas rates. Crum (1981) maintains that at below 10 MMscfd (300 MSm3/d), turboexpanders offer less economic advantage and they lose efficiency below 5 MMScfd (150 MSm3/d).
• Low ethane recovery. For ethane recoveries of 10 to 30%, J-T expansion may be sufficient.
• Variable flow rates. J-T is insensitive to widely varying flow rates, whereas turboexpanders lose efficiency when operating off of design rates.

Crum (1981) also points out that J-T plants are much simpler than turboexpander plants because J-T plants have no need for seal gas and lubricating oil systems. However, because of the inefficiency of J-T valves compared with turboexpanders, if any inlet compression is required, more is required with J-T expansion to obtain the same amount of refrigeration. The Engineering Data Book, suggests that use of J-T expansion for limited ethane recovery requires inlet pressures around 1,000 psi (70 bar).
7.3.2.2 Lean Oil Absorption
Early gas processing plants used lean oil absorbers to strip NGL from natural gas, and the process is still used in about 70 gas plants today.
To improve recoveries, later plants used external refrigeration to cool the feed gas and lean oil. Figure 7.9 shows a representative schematic of a propane refrigerated lean oil system. The process involves three steps:
1. Absorption. An absorber contacts a lean oil to absorb C2+ plus from raw natural gas.
2. Stabilization. The rich oil demethanizer (ROD) strips methane and lighter components from the rich oil.
3. Separation. The still separates the recovered NGL components as product from the rich oil, and the lean oil then returns to the absorber.
Gas from the ROD is either blended with the exiting gas stream or used for fuel. Original systems used lean oil with molar mass of 150 to 200, but refrigerated systems use molar masses of 100 to 130. If no refrigeration is used, and assuming the absorber runs at about 100°F (38°C), over 75% of the butanes and essentially all of the C5+ fraction are recovered. Using high solvent rates makes possible the recovery of 50% of the ethane and essentially all of the propane and heavier components. With propane refrigeration, typically over 97% of the propane is recovered and up to 50% of the ethane.
However, the process is energy intensive and relies on numerous heat exchangers to reduce the energy load. For gas processing, the whole process can be simplified by elimination of the lean oil and use of external refrigeration, as discussed in the previous section. However, many refrigerated lean oil absorption plants remain in operation today with capacities of 1,000 MMscfd (30 MMSm3/d) or more. One use for lean oil absorbers today is in capturing fugitive hydrocarbons from air streams because refrigeration is unnecessary. Many non gas related industries use this process for pollution control.

7.3.3 High Ethane Recovery
The above processes provided limited recovery of ethane. To obtain 80 to 90% or more ethane recovery requires separation temperatures well below what is obtainable by use of propane refrigeration alone. In principle, direct-refrigeration processes could be used by cascading propane cooling with ethane or ethylene refrigeration or by use of a mixed refrigerant that contains methane, ethane, and propane. The primary motivation for use of only direct refrigeration would be low inlet gas pressures. If significant inlet compression is required to produce refrigeration by expansion, then cascade or mixed-refrigeration cooling, with or without expansion, may be attractive. No matter which option is used, obtaining high ethane recoveries from low inlet-pressure feed streams requires substantial compression, of either the feed stream, the refrigerants, or both.
With recovery of a high ethane fraction, sales gas specifications must be considered.
Recovery of too much ethane could reduce the heating value below contract limits.
Figure 7-10 shows a simplified conventional expander plant schematic. It consists of a gas−gas heat exchanger with five gas streams that enter at different temperatures, cold separator, turboexpander, and demethanizer. Although the flow sheet shown is schematically simple, in practice most actual designs replace the single exchanger with a more complex and efficient combination of exchangers.
The inlet gas stream makes several passes through the gas−gas exchanger before going to the cold separator, where the vapor expands through a turboexpander. Liquid from the cold separator is flashed through a J-T valve and fed to the middle of the demethanizer. The incoming gas provides reboiler heat at the bottom, and then is cooled further in a second reboiler midway up the column.
A J-T valve is always installed parallel to the turboexpander. This configuration helps in plant start-up and in handling excess gas flow. It also is used if the turboexpander goes down.
The maximum ethane recovery with the conventional turboexpander configuration is about 80%. Also, the cold separator may be near the critical temperature and pressure of the mixture, which can make the process unstable. Carbon dioxide freezing out can also be a problem. Improvement of C2+ recovery requires reduction of ethane losses in the top of the demethanizer by addition of reflux. The Engineering Data Book discusses a number of configurations used. One that can provide up to 98% recovery, called the cold residue recycle (CRR) process, is shown in Figure 7-11, which gives the maximum ethane recovery with regards to compression requirements of all commonly used processes. It has the added advantage that it can reject ethane and still maximize C3+ recovery if desired. In this variation, the cold separator runs at a warmer temperature to avoid the critical point problem. The vapor from the separator splits into two streams. Part goes to the turboexpander and the balance goes through two overhead exchangers, where it is condensed to provide liquid reflux to the column. Turboexpander output then enters further down the column. In addition, part of the overhead is compressed and cooled to provide additional reflux.


Fig. 7.9 Refrigerated lean oil absorption process.


Fig. 7-10 Schematic of conventional turboexpander process with no recycle to demethanizer. Note that the one heat exchanger represents a network of exchangers.


Fig.7-11 Cold-residue recycle process for maximizing ethane recovery. All valves in figure are J-T expander valves but are unlabeled for figure clarity and the large heat exchanger represents a network of exchangers.

 
Basics of Natural Gas Field Processing
References.
1- Fundamentals of oil and Gas Processing – Yasser Kassem – Amazon.com Publication. 2018.
2- Gas Processors Suppliers Association GPSA Engineering Data Book 11th, 12th, & 15th Editions. Tusla, OK.
3- Arnold, K. and Stewart, M., Surface prod operations V1_ 2E, Surface prod operations V2_ 2E, & Surface prod operations V1_ 3E, Gulf Publishing Co., Richardson, TX.
4- Fundamentals of Natural Gas Processing - Arthur J. Kidnay- William R. Parrish- 2006 by Taylor and Francis Group, LLC.
5- Design, Operation and Maintenance of Gas Plants- Campbell, John M. 2003. BP EXPLORATION COMPANY (COLUMBIA) LTD.
6- Oil Field Processing of Petroleum- Volume 1 – Natural Gas. Franci S. Manning. Penn Well Publishing Co. 1991.
7- Abdel-Aal, H. K., Surface Petroleum Operations, Saudi Publishing & Distributing House, Jeddah, 1998.
8- H.K. Abdel-Aal and Mohamed Aggour, Petroleum and gas field processing, 2003 by Marcel Dekker, Inc.
9- Crude-Oil-Treating-Systems-Design-Manual-Sivalls-Inc.
10- API Spec. 12J (Specification of Oil and Gas Separator) 7th.ed. Oct. 1998.
11- Standard Handbook of Petroleum and Natural Gas 2nd ed. William C. Lyons, Ph.D., P.E. Gary J. Plisga, B.S. 2005, Elsevier Inc.
12- Gas Pipeline Hydraulics, E. Shashi Menon, P.E. PDH Engineering course material.
13- Chilingarian, G. V., Robertson, J. O., Jr., and Kumar, S., Surface Operations in Petroleum Production, I & 2, 1987, Elsevier Science, Amsterdam.
14- The Chemistry and Technology of Petroleum, James G. Speight
15- Flow Management for Engineers and Scientists, Nicholas P. Cheremisinoff and Paul N. Cheremisinoff.
16- Campbell, John M., ‘‘Gas Conditioning and Processing,’’ Vol. 2, published by Campbell Petroleum Series, Norman, Oklahoma, 1976.
17- Hybrid Systems- Combining Technologies Leads to More Efficient Gas Conditioning - William Echt, UOP LLC- UOP technical publications.
18- HANDBOOK OF NATURAL GAS TRANSMISSION AND PROCESSING - Saeid Mokhatab, William A. Poe & James G. Speight - 2006, Elsevier Inc.
19- NATURAL GAS HYDRATES IN FLOW ASSURANCE - 2011 Dendy Sloan, Carolyn Ann Koh, Amadeu K. Sum, Norman D. McMullen, George Shoup, Adam L. Ballard, and Thierry Palermo. Published by Elsevier Inc.

[bebopetro47]

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[yasserkassem]

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