Fundamentals of Oil and Gas Processing Book "full text"

yasserkassem · Post by **yasserkassem** » Tue Mar 17, 2020 9:28 pm

Chapter #1

Chapter 1 10
Basics of Oil and Gas Treatment 10
1.1 Introduction 10
1.2 Hydrocarbon preparation 10
1.3: Physical properties of Hydrocarbon Gases 11
1.3.1: Hydrocarbon gases 11
1.3.2: Molecular weight and apparent molecular weight 11
1.3.3: Apparent molecular weight of gas mixture 12
1.3.4: Gas Specific Gravity and Density 13
1.3.6: Compressibility Factor (z) 14
1.3.7: Gas density at any condition of Pressure and temperature 18
1.3.8: Gas volume at any condition of Pressure and temperature 19
1.3.9: Velocity of gas, (ft/s) 20
1.3.10: Average pipeline pressure 21
1.3.11: Viscosity of gases 22
1.3.12: The heating value of gases 22
1.4: properties of Hydrocarbon Liquids (Crude Oil) 23
1.4.1: Introduction 23
1.4.2: Crude oil Density and gravity 24
1.4.3: Crude oil Viscosity. 25
1.4.4: Oil-Water Mixture Viscosity 25
1.5: Phase Behavior 27
1.5.1: Introduction 27
1.5.2 System Components 27
1.5.3: Single-Component Systems 28
1.5.4: Multicomponent Systems 31
1.5.5: Prediction of phase envelope 32
1.6: Types of Fluid Flow 42
1.6.1: Reynolds Number 42
Chapter 2 43
Two-phase Oil and Gas Separation 43
2.1 Introduction 43
2.2 Phase Equilibrium 43
2.3: Separation process: 43
2.4: Principles of Physical Separation: 44
2.5: Gravity Separation: 44
2.6: Factors Affecting Separation 46
2.7: Separator categories and nomenclature: 47
2.8: Functional Sections of a Gas-Liquid Separator 47
2.8.1: Inlet Diverter Section 48
2.8.2: Liquid Collection Section 48
2.8.3: Gravity Settling Section 48
2.8.4: Mist Extractor Section 49
2.9: Separator Configurations 49
2.10: Types of Separators 50
2.10.1: Vertical Separators 50
2.10.2: Horizontal Separators 52
2.10.3: Double-Barrel Horizontal Separators 53
2.10.4: Horizontal Separator with a “Boot” or “Water Pot” 54
2.10.5: Filter Separators 54
2.10.6: Scrubbers 56
2.10.7: Slug Catchers 56
2.11: Selection Considerations 57
2.12: Internal Vessel Components 59
2.12.1: Inlet Diverters 59
2.12.2: Wave Breakers 62
2.12.3: Defoaming Plates 62
2.12.4: Vortex Breaker 63
2.12.5: Stilling Well 64
2.12.6: Sand Jets and Drains 64
2.12.7: Mist Extractors 65
2.13: Control Components of Gas–Oil Separators 76
2.14.1: Foamy Crude 77
2.14.2: Paraffin 78
2.14.3: Sand 78
2.14.4: Gas Blowby 78
2.14.5: Liquid Carryover 79
2.14.6: Liquid Slugs 79
2.15: Stage Separation 80
2.15.1: Initial Separation Pressure 80
2.15.2: Stage Separation 81
2.15.3: Selection of Stages 83
2.15.4: Fields with Different Flowing Tubing Pressures 83
2.15.5: Determining Separator Operating Pressures 84
2.15.6: Two-Phase vs. Three-Phase Separators 85
2.16: Separator calculation basics. 85
2.16.1: Liquid Handling and Liquid Retention Time 85
2.16.2: Gas retention time 86
2.16.3: Gas velocity 86
2.16.4: Liquid Re-entrainment 87
2.16.5: Droplet Size (Liquid in gas phase) 88
2.17: Design Principles and sizing of Oil-gas Separator 88
2.17.1: First method Design Theory 89
2.17.1.12: Slenderness Ratio 95
2.17.2: Second method Design Theory 100
Chapter 3 107
Three-phase Oil and Gas Separation 107
3.1: Introduction 107
3.2: three phase separation equipment’s 108
3.2.1: Horizontal Separators 108
3.2.2: Free-Water Knockout 111
3.2.3: Horizontal Three-Phase Separator with a Liquid “Boot” 111
3.2.4: Vertical Separators 112
3.2.5: Selection Considerations 114
3.3: Internal Vessel components 115
3.3.1: Coalescing Plates 117
3.4: Operating Problems 118
3.4.1: Emulsions 118
3.5: Three-Phase Separator Design Theory 118
3.5.1: Gas Separation 118
3.5.2: Oil–Water Settling 118
3.5.3: Water Droplet Size in Oil 118
3.5.4: Oil Droplet Size in Water 119
3.5.5: Retention Time 119
3.6: Separator Design (first method) 121
3.6.1: Horizontal Three-phase Separator Sizing—Half-Full 121
3.6.1.2: Retention Time Constraint 121
3.6.1.3: Settling Water Droplets from Oil Phase 122
3.6.1.4: Separating Oil Droplets from Water Phase 123
3.6.2: Vertical Separators’ Sizing 124
3.6.2.1: Gas Capacity Constraint 125
3.6.2.3: Settling Oil from Water Phase Constraint 125
3.7: Separator Design (second method) 131
Chapter 4 134
Crude oil dehydration 134
4.1: Introduction 134
4.2: Emulsion 134
4.2.1 Energy of Agitation 135
4.2.2 Emulsifying Agents 136
4.2.3: Stability of oil water emulsion 137
4.2.4: Emulsion Treating Theory 139
4.2.5: Demulsifiers 140
4.3: Crude oil treating systems 143
4.3.1: Free-Water Knockouts 143
4.3.2: Gunbarrel tanks with internal and external gas boots 144
4.3.3: Heaters 146
4.4: Emulsion Treating Methods 164
4.4.1: General Considerations 164
4.4.2: Chemical Addition 165
4.5: Heat Required 174
4.5.1: Heat duty 174
4.5.2: Heat Loss 174
4.5.3: Fire Tube Heat Flux 175
4.5.4: Firetube Heat Density 175
4.6: Treater Equipment Sizing 175
4.6.2: Design Procedure 178
4.7: Practical Considerations 184
4.7.1: Gunbarrels with Internal/External Gas Boot 184
4.7.2: Heater-Treaters 184
4.7.3: Electrostatic Heater-Treaters 184
Chapter 5 185
Crude Oil Desalting 185
5.1: Introduction 185
5.1.1: Salt Content 185
5.1.2: Desalting Process 186
5.2: Equipment Description 186
5.2.1: Desalters 186
5.2.2: Mixing Equipment 186
5.3: Process Description 188
5.3.1: Single-Stage Desalting 189
5.3.2: Two-Stage Desalting 189
5.4: Electrostatic Desalting Voltage 189
5.5: Operating Parameters Effects 191
5.6: Design Consideration 191
5.7: Troubleshooting 192
Chapter 6 193
Crude Oil Stabilization and Sweetening 193
6.1: Introduction 193
6-1-1: Crude oil treatment steps 193
6.2: Process Schemes 194
6.2.1: Multi-Stage Separation 194
6.2.2: Oil Heater-Treaters 194
6.2.3: Liquid Hydrocarbon Stabilizer 195
6.2.4: Cold-Feed Stabilizer 197
6.2.5: Stabilizer with Reflux 197
6.3: Stabilization Equipment 199
6.3.1: Stabilizer Tower 199
6.4: Stabilizer Design 205
6.5: Crude Oil Sweetening 206
6.6.1: Stage vaporization with stripping gas. 206
6.6.2: Trayed stabilization with stripping gas. 207
6.6.3: Reboiled trayed stabilization. 208
Chapter 7 209
Fluid Measurements 209
7.1: Gas Measurement 209
7.1.1: Orifice-Meter Measurement 209
7.1.1.5: Meter Tubes 213
7.1.2: Ultrasonic Measurement 220
7.2: Liquid Measurements 221
7.2.1: Volumetric Measurement Meters (Orifice Meters) 221
7.2.2: Turbine Meters 223
7.2.3: Positive Displacement Meters 224
7.2.4: Turbine and Positive Displacement Meter Selection 224
7.2.5: Mass Measurement Meters 225
Chapter 8 228
Instrumentation and Control 228
8.1: Introduction 228
8.2: Type Selection and Identification 228
8.2.1: Pneumatic Power Supplies 228
8.2.2: Electronic Power Supplies 229
8.3: Sensing Devices 230
8.3.1: Pressure Sensors 230
8.3.1.3: Bellows (Fig. 8-3) 230
8.3.2: Level Sensors 232
8.3.3: Temperature Sensors 237
8.3.4: Flow Sensors 239
8.4: Signal Transmitters 241
8.4.1: Pneumatic Transmitters 241
8.4.2: Electronic Transmitters 241
8.5: Signal Converters 241
8.5.1: Pneumatic-to-electronic (P/I) 242
8.5.2: Electronic-to-pneumatic (I/P) 242
8.5.3: Isolators 242
8.5.4: Electric signal converters 242
8.5.5: Frequency converters 242
8.6: Recorders and Indicators 242
8.6.1: Recorders 242
8.6.2: Indicators 242
8.7: Control Concepts 243
8.7.1: Control Loops 243
8.8: Control Modes and Controllers 245
8.8.1: Two-Position (on-off) Controllers 245
8.8.2: Proportional Control Mode 245
8.9: Control Valves 246
8.9.1: Control-Valve Bodies 247
8.9.2: Control-Valve Actuators 248
8.9.3: Flow Characteristics and Valve Selection 249
8.9.4: Fundamentals of Control Valve Sizing 250
Chapter 9 256
Process Relief Systems 256
9.1: Introduction 256
9.2: Relief Device Design and Requirements: 256
9.2.1: Blocked Discharge 257
9.2.2: Fire Exposure 257
9.2.3: Tube Rupture 257
9.2.4: Control Valve Failure 257
9.2.5: Thermal Expansion 257
9.2.6: Utility Failure 257
9.3: General discussion 258
9.4: Special Relief System Considerations 260
9.4.1: Pumps and storage equipment 260
9.4.2: Low Temperature Flaring 260
9.5: Relieving Devices 260
9.5.1: Conventional Relief Valves 260
9.5.2: Balanced Relief Valves 262
9.5.3: Pilot Operated Relief Valves 262
9.5.4: Resilient Seat Relief Valves 264
9.5.5: Rupture Disk 265
References. 267

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Chapter 1

Basics of Oil and Gas Treatment
1.1 Introduction
Oil and gas wells produce a mixture of hydrocarbon gas, condensate or oil, salt water, other gases, including nitrogen, carbon dioxide (CO2), and possibly hydrogen sulfide (H2S), and solids, including sand from the reservoir, dirt, scale, and corrosion products from the tubing. These mixtures are very difficult to handle, meter, or transport. In addition to the difficulty, it is also unsafe and uneconomical to ship or to transport these mixtures to refineries and gas plants for processing. Further, hydrocarbon shipping tankers, oil refineries, and gas plants require certain specifications for the fluids that each receive. Also, environmental constraints exist for the safe and acceptable handling of hydrocarbon fluids and disposal of produced salt water. It is therefore necessary to process the produced fluids in the field to yield products that meet the specifications set by the customer and are safe to handle.
1.2 Hydrocarbon preparation
The goal is to produce oil that meets the purchaser’s specifications that define the maximum allowable amounts of water, salt, and sulfur. In addition to the maximum allowable value of Reid vapor pressure and maximum allowable pour point temperature.
Similarly, the gas must be processed to meet purchaser’s water vapor maximum allowable content (Water dew point), hydrocarbon dew point specifications to limit condensation during transportation, in addition to the maximum allowable content of CO2, H2S, O2, Total Sulfur, Mercaptan, Mercury, and maximum gross heating value.
The produced water must meet the regulatory requirements for disposal in the ocean if the wells are offshore, or to meet reservoir requirements for injection into an underground reservoir to avoid plugging the reservoir.
The specifications for the above requirements may include maximum oil in water content, total suspended solids to avoid formation plugging, bacteria counts, toxicity in case of offshore disposal, and oxygen content. Before discussing the industry or the technology of oil and gas processing it is best to define the characteristic, physical properties and main chemical composition of oil and gas produced.
Figures 1-1 and 1-2, illustrates gas-oil separation plant, and oil flow diagram.

Fig.1-1 .Gas Oil Separation plant function.

Notes.
Separator may be a slug catcher, free water knock out drum, two phase separator, or gun barrel.
A dehydrator may be a heater treater, separator, or settling tank.
Heat is added upstream or downstream separator depending on crude oil temperature and gas oil ratio.
Crude oil stabilization is usually performed in separation step or during heat addition.
Crude oil sweetening is usually performed upstream or downstream heater treater.
Gas and water are separated and undergoes further treatment processes not in the scope of this book.

Fig.1-2. Crude oil flow Diagram

1.3: Physical properties of Hydrocarbon Gases
1.3.1: Hydrocarbon gases
Most of compounds in crude oil and natural gas consist of molecules made up of hydrogen and carbon, therefore these types of compounds are called hydrocarbon.
The smallest hydrocarbon molecule is Methane (CH4) which consists of one atom of Carbon and four atoms of hydrogen. It may be abbreviated as C1 since it consisted from only one carbon atom. Next compound is Ethane (C2H6) abbreviated as C2, and so on Propane (C3H8), Butane (C4H10)...etc.
Hydrocarbon gases are C1:C4), with the increase of carbon number, liquid volatile hydrocarbon is found (e.g. Pentane C5 is the first liquid hydrocarbon at standard conditions).

1.3.2: Molecular weight and apparent molecular weight
The molecular weight of a compound is the sum of the atomic weight of the various atoms making up that compound. The Mole is the unit of measurements for the amount of substance, the number of moles is defined as follows:

Mole = Weight/(Molecular weight) (Eq. 1-1)

Expressed as n = m/M (Eq. 1-2)
or, in units as, lb-mole = lb/(lb/lb-mole) (Eq. 1-3)

Table 1-1 Physical constants of light hydrocarbons and some inorganic gases. Adapted from GPSA, Engineering Hand Book.

Example 1.1: Methane molecule consists of one carbon atom with atomic weight =12 and 4 hydrogen atoms with atomic weight = 1 each. Molecular weight for Methane (CH4) = (1 × 12) + (4 × 1) = 16 lb/lb-mole. Similarly, Ethane (C2H6) molecular weight = (2 × 12) + (6 × 1) = 30 lb/lb-mole.

Hydrocarbon up to four carbon atoms are gases at room temperature and atmospheric pressure. Reducing the gas temperature and/or increasing the pressure will condense the hydrocarbon gas to a liquid phase. By the increase of carbon atoms in hydrocarbon molecules, consequently the increase in molecular weight, the boiling point increases and a solid hydrocarbon is found at high molecular weight.
Physical constants of light hydrocarbon and some inorganic gases are listed in Table 1-1.

1.3.3: Apparent molecular weight of gas mixture
For compounds, the term molecular weight is used, while, for hydrocarbon mixture the term apparent molecular weight is commonly used. Apparent molecular weight is defined as the sum of the products of the mole fractions of each component times the molecular weight of that component. As shown in Eq. 1-4
MW= ∑▒ Yi (MW)i (Eq. 1-4)
where
yi =molecular fraction of ith component,
MWi =molecular weight of ith component,
Ʃyi =1.

Example 1.2: Determine the apparent molecular weight for the gas mixture in Table 1-2:

Table 1-2 Gas mixture for Example 1-2

Solution: Using Table 1-1 & Equation 1-4
MW= ∑▒ Yi (MW)i
MW = (Mole Fraction of component 1 × MW of component 1) + (Mole Fraction of component 2 × MW of component 2) + (Mole Fraction of component 3 × MW of component 3) +…etc.
The following table can be made:

Table 1-3 Solution of Example 1-2
The apparent molecular weight is 21.36

1.3.4: Gas Specific Gravity and Density
The density of a gas is defined as the mass per unit volume as follows
Density = mass / volume (Eq. 1-5)

The specific gravity of a gas is the ratio of the density of the gas to the density of air at standard conditions of temperature and pressure.
S = (ρ(gas))/(ρ(air)) (Eq. 1-6)
Where
ρ(gas) ρg = density of gas
ρ(air) ρair = density of air

Both densities must be computed at the same pressure and temperature, usually at standard conditions.
It may be related to the molecular weight by Equation 1-7
S = (MW(gas))/(MW(air)) (Eq. 1-7)
Since molecular weight of air is 28.96 (table 1-1)
Specific gravity of gas S = (MW(gas))/28.96 (Eq. 1-8)

Example 1-3: Determine the specific gravity of the gas mixture in example 1-2.
Solution:
Apparent molecular weight of gas mixture is 21.36
Gas specific gravity = 21.36/28.96 = 0.7376

Since the gas is a compressible fluid, its density varies with temperature and pressure, calculating the gas density at a certain pressure and temperature will be explained after discussing the general gas law and gas compressibility factor.

1.3.5: General Gas Law
The general (Ideal) Gas equation, or the Perfect Gas Equation, is stated as follows:

PV = nRT (Eq. 1-9)

Where
P = gas pressure, psia
V = gas volume, ft3
n = number of lb moles of gas (mass/molecular weight)
R = universal gas constant, psia ft3/lb mole OR
T = gas temperature, OR (OR = 460 + OF)
The universal gas constant R is equal to 10.73 psia ft3/lb mole OR in field units.

Equation (1-9) is valid up to pressures of about 60 psia. As pressure increases above this level, its accuracy becomes less and the system should be considered a non-ideal gas equation of state.
PV = znRT (Eq. 1-10)

Where
z = gas compressibility factor.

1.3.6: Compressibility Factor (z)
The Compressibility factor, Z is a dimensionless parameter less than 1.00 that represents the deviation of a real gas from an ideal gas. Hence it is also referred to as the gas deviation factor. At low pressures and temperatures Z is nearly equal to 1.00 whereas at higher pressures and temperatures it may range between 0.75 and 0.90. The actual value of Z at any temperature and pressure must be calculated taking into account the composition of the gas and its critical temperature and pressure. Several graphical and analytical methods are available to calculate Z. Among these, the Standing-Katz, and CNGA methods are quite popular. The critical temperature and the critical pressure of a gas are important parameters that affect the compressibility factor and are defined as follows.
The critical temperature of a pure gas is that temperature above which the gas cannot be compressed into a liquid, however much the pressure. The critical pressure is the minimum pressure required at the critical temperature of the gas to compress it into a liquid.
As an example, consider pure methane gas with a critical temperature of 343 0R and critical pressure of 666 psia (Table 1-1).
The reduced temperature of a gas is defined as the ratio of the gas temperature to its critical temperature, both being expressed in absolute units (0R). It is therefore a dimensionless number.
Similarly, the reduced pressure is a dimensionless number defined as the ratio of the absolute pressure of gas to its critical pressure.
Therefore we can state the following:
Tr = T/Tc (Eq. 1-11)
Pr = P/Pc (Eq. 1-12)

Where
P = pressure of gas, psia
T = temperature of gas, 0R
Tr = reduced temperature, dimensionless
Pr = reduced pressure, dimensionless
Tc = critical temperature, 0R
Pc = critical pressure, psia

Example1-4: Using the preceding equations, the reduced temperature and reduced pressure of a sample of methane gas at 70 0F and 1200 psia pressure can be calculated as follows

Tr = (70 +460) / 343 =1.5
Pr = 1200/666 = 1.8

For natural gas mixtures, the terms pseudo-critical temperature and pseudo-critical pressure are used. The calculation methodology will be explained shortly. Similarly we can calculate the pseudo-reduced temperature and pseudo-reduced pressure of a natural gas mixture, knowing its pseudo-critical temperature and pseudo-critical pressure.
The Standing-Katz chart, Fig. 1.3 can be used to determine the compressibility factor of a gas at any temperature and pressure, once the reduced pressure and temperature are calculated knowing the critical properties.
Pseudo-critical properties allow one to evaluate gas mixtures. Equations (1-13) and (1-14) can be used to calculate the pseudo-critical properties for gas mixtures:

P’c = Ʃ yi Pci (Eq. 1-13)

T’c = Ʃ yi Tci (Eq. 1-14)

where
P’c =pseudo-critical pressure,
T’c =pseudo-critical temperature,
Pci =critical pressure at component i, psia
Tci =critical temperature at component i, 0R
Yi =mole fraction of each component in the mixture,
Ʃ yi =1.

Example 1-5: Calculate the Compressibility factor for the following Gas mixture at 1000F and 800 psig:

Table 1-4 for Example 1-5.

Using Equation 1-11 and 1-12
T`r = (100+460)/464.5 =1.2
P`r = (800+14.7)/659.4 = 1.23
From fig.1-3. Compressibility factor is approximately, z= 0.72

Calculating the compressibility factor for example 1-4, of the gas at 70 0F and 1200 psia, using Standing-Katz chart, fig. 1-3. Z = 0.83 approximately. For ) Tr = 1.5 , Pr = 1.

.

Another analytical method of calculating the compressibility factor of a gas is using the CNGA equation as follows:

(Eq. 1-15)
Where
Pavg = Gas pressure, psig. [psig = (psia - 14.7)]
Tf = Gas temperature, 0R
G = Gas gravity (air = 1.00)
The CNGA equation for compressibility factor is valid when the average gas pressure Pavg is greater than 100 psig. For pressures less than 100 psig, compressibility factor is taken as 1.00. It must be noted that the pressure used in the CNGA equation is the gauge pressure, not the absolute pressure.

Example 1-6: Calculate the compressibility factor of a sample of natural gas (gravity = 0.6) at 80 0F and 1000 psig using the CNGA equation.
Solution:
From the Eq. (1.15), the compressibility factor is

The CNGA method of calculating the compressibility, though approximate, is accurate enough for most gas pipeline hydraulics work and process calculations.

Figure 1-3 Compressibility Factor For lean sweet natural gas (Surface Production Operations).

1.3.7: Gas density at any condition of Pressure and temperature
Once the molecular weight of the gas is known, the density of a gas at any condition of temperature and pressure is given as:

ρg= ((MW)P)/RTZ lb/ft3

Since R=10.73, then
ρg= 0.093 ((MW)P)/TZ lb/ft3 (Eq. 1-16)
where
ρg = density of gas, lb/ft3,
P =pressure, psia,
T =temperature, 0R,
Z =gas compressibility factor,
MW=gas molecular weight.

Example 1-7: Calculate the pseudo-critical temperature and pressure for the natural gas stream composition given in example 1-2, calculate the compressibility factor, and gas density at 600 psia and 1000F.
Solution:

Table 1-5 solution of Example 1-7.
From the table MW= 21.36
T`c = 451.5 0R
P`c = 667 psia

From Eq. (1-11) and Eq. (1-12)
Tr = T/T`c = (100+460)/451.5 = 1.24

Pr = P/P`c = 600/667 = 0.9

Compressibility factor z could be calculated from figure 1-3, or from Eq. (1-15)
Value from figure, z = 0.83
From Equation 1-15 z = 0.87
For our further calculation we will use the calculated z value [Eq. (1-15)]
Using eq. (1-16) density of gas
ρg = 0.093 ((21.36)600)/(560 ×0.83) = 2.56 lb/ft3
Comparing ρg at standard condition (z=1)
ρg at standard condition = 0.093 (21.36)14.7/(520 ×1) = 0.056 lb/ft3
We can conclude that density increases with pressure while the volume decreases.

1.3.8: Gas volume at any condition of Pressure and temperature
Volume of a gas is the space occupied by the gas. Gases fill the container that houses the gas. The volume of a gas generally varies with temperature and pressure.
Volume of a gas is measured in cubic feet (ft3).
Gas volume are commonly referred to in "standard" or "normal" units.
Standard conditions commonly refers to gas volumes measured at: 60°F and 14.696 psia
The Gas Processors Association (GPA) SI standard molar volume conditions is 379.49 std ft3/lb mol at 60°F, 14.696 psia.
Therefore, each mole (n) contains about 379.5 cubic feet of gas (ft3)at standard conditions.
Therefore, by knowing the values of mass and density at certain pressure and temperature, the volume occupied by gas can be calculated.

Example 1-8: Calculate the volume of a 10 lb mass of gas (Gravity = 0.6) at 500 psig and 80 0F, assuming the compressibility factor as 0.895. The molecular weight of air may be taken as 29 and the base pressure is 14.7 psia.
Solution:
The molecular weight of the gas (MW) = 0.6 x 29 = 17.4
Pressure =500+14.7 = 514.7 psia
Temperature = 80+460 = 540 0R
Compressibility factor z= 0.895
The number of lb moles n is calculated using Eq. (1-2). n=m/(MW)
n = 10/17.4
Therefore n= 0.5747 lb mole
Using the real gas Eq. (1-10), PV=nzRT
(514.7) V = 0.895 x 0.5747 x 10.73 x 540. Therefore, V = 5.79 ft3

Example 1-9: Calculate the volume of 1 lb mole of the natural gas stream given in the previous example at 1200F and 1500 psia (compressibility factor Z = 0.811).
Solution:
Using Eq.(1-10), PV = nzRT
V= 0.811 x 1 x 10.73 x (120+460)/1500. V = 3.37 ft3

Example 1-10: One thousand cubic feet of methane is to be compressed from 60°F and atmospheric pressure to 500 psig and a temperature of 50°F. What volume will it occupy at these conditions?
Solution:
Moles CH4 (n) = 1000 / 379.5 = 2.64
At final conditions, (Compressibility factor z must be calculated), from equations 1-11 and1-12
Tr = (460 + 50) / 344 = 1.88
Pr = (500 + 14.7) / 673 = 0.765
From Figure 1-3, Z = 0.94
From eqn. 1-10, PV = nzRT
V = ft3
Example 1-11: One pound-mole of C3 H8 (44 lb) is held in a container having a capacity of 31.2 cu ft. The temperature is 280°F. "What is the pressure?
Solution:
Volume = V = 31.2 ft3
A Trial-and-error solution is necessary because the compressibility factor Z is a function of the unknown pressure. Assume Z = 0.9.
Using Eq. 1-10, PV = nzRT
P ×31.2 = 0.9 × 1.0 × 10.7 × (460 + 280)
P = 229 psia
From table 1-1, eqns. 1-11 and 1-12
Pr = 229 / 616 = 0.37,
Tc = 665ºR
Tr = (460 + 280) / 665 = 1.113
According to Figure 1.3, the value of Z should be about 0.915 rather than 0.9. Thus, recalculate using eq. 1-10, the pressure is 232 rather than 229 psia.

Example 1-12: Calculate the volume of gas (MW=20) will occupy a vessel with diameter 24 in, and 6 ft. length. At pressure 200 psia and temperature 100 0F. (Assume compressibility factor z=0.9), and what will be the volume of gas at 14.7 psia and 60 0F.
Then calculate gas density and mass inside the container at pressure 200 psia and temperature 100 0F.
Volume of vessel = π L r2
V = 3.14 × 6 × (24)2/ (2 ×12)2 ft3
V = 18.8 ft3.
(We divided by 2 to get r from the diameter, and divided by 12 to convert from in. to ft.)
T = 460 + 100 = 560 0R
Using Eq. 1-10, PV=nZRT
n = 18.8 × 200 / (0.9 × 10.73 × 560)
n = 0.7 lb. moles. (Remember gas volume ft3 = 379.5 x n)
Volume of gas at 200 psia and 100 0F= 0.7 * 379.5 = 266 ft3
n of Gas at 14.7 psia and 60 0F ( z=1) = 18.8 × 14.7 / (1 × 10.73 × 520)
n = 0.0495 lb. moles
Volume of gas at 14.7 psia and 60 0F = 0.0495 * 379.5 = 18.8 ft3
From the previous example 1-12, the gas volume will equal to the container volume at standard conditions (14.7 psia and 60 0F).
Gas density is calculated using Eq. 1-16
ρg = 0.093 ((MW)P)/TZ lb/ft3
Density of gas ρg = 0.093 × 20 × 200 / (0.9 × 560) = 0.738 lb/ft3
Mass of gas inside the vessel = Volume × density = 0.738 × 265 = 196 lb mass

1.3.9: Velocity of gas, (ft/s)
The velocity of gas equal the volume flow rate (ft3) per second divided by flow area (ft2).

Example 1-13: Calculate the gas velocity for gas flow rate 100 MMscfd through 24 in. internal diameter gas pipe, the gas specific gravity is 0.7, pressure 500 psia, Temperature 100 0F, and assume compressibility factor 0.85.
Solution: Using Eq. 1-10, PV=nzRT, and remember that n= V (ft3)/379.5).
n = 100 × 106/379.5
Gas volume at operating conditions V= 100 × 106 × 0.85 × 10.73 × 560 / (379.5 × 500)
= 2,695,000 ft3/day
Gas flow rate cubic foot per second = 2,695,000 / (24×60×60) = 31.2 ft3/sec
Area of flow = π r2 = 3.14 × 12 × 12 / (144) = 3.14 ft2
(144 to convert r2 from in. to ft2.)
Velocity of gas will be 31.2/3.14 = 9.9 ft/s
The gas velocity may be calculated directly from the following equation:
Velocity = 6 ZTQ/(100,000×Pd2) ft/s. Eq 1-17
Where Q = Flow rate scfd, d = diameter in inches.

The maximum recommended velocity of dry gas in pipes is 100 ft/s, (60 ft/s for wet gas), and to be less than the erosional velocity which is defined as:
Erosional velocity: The erosional velocity represents the upper limit of gas velocity in a pipeline. As the gas velocity increases, vibration and noise result. Higher velocities also cause erosion of the pipe wall over a long time period. The erosional velocity Vmax may be calculated approximately as follows:

Vmax = 100 √(2&ZRT/29GP) Eq 1-18

Where G= gas sp. Gt (air=1), P = pressure psia
For Example 1-12, the erosional velocity Vmax is:
Vmax = 100 √(2&0.85×10.73×560/(29×0.7× 500)) Vmax = 70.9 ft/s.

1.3.10: Average pipeline pressure
The gas compressibility factor Z used in the General Flow equation is based upon the flowing temperature and the average pipe pressure. The average pressure may be approximated as the arithmetic average
Pavg = (P1+P2)/2 of the upstream and downstream pressures P1 and P2. However, a more accurate average pipe pressure is usually calculated as follows
Pavg = 2/3 (P1+P2 - (P1× P2)/(P1+ P2)) Eq 1-19
Where
P1, P2, Pavg = pressure, psia

Example 1-14: A natural gas pipeline with internal diameter 19 in. transports natural gas (Sp. Gr.= 0.65) at a flow rate of 200 MMscfd. Calculate the gas velocity at inlet and outlet of the pipe, assuming isothermal flow. The inlet temperature of 70 0F, inlet pressure is 1200 psig, and outlet pressure is 900 psig. Use compressibility factor of 0.95. Also, calculate the erosional velocity for this pipeline.
Solution:
Using Eq. 1-17, the gas velocity at inlet of the pipe:
Velocity = 6 × 0.95× 530×200,000,000/(100,000×1214.7×192) ft/s.
Velocity = 13.8 ft/s.
The gas velocity at outlet of the pipe:
Velocity = 6 × 0.95× 530×200,000,000/(100,000×914.7×192) ft/s.
Velocity = 18.3 ft/s.
Finally, the erosional velocity can be calculated using Eq. 1-18
Vmax = 100 √(2&0.95 ×10.73×530/29×0.65×1214.7)
Vmax = 48.6 ft/s.

The above example may be solved by calculating the gas density at inlet and outlet of the pipe, then calculating the operational flow rate, divide it by pipe cross sectional area to get the velocity as follows:
Gas molecular weight = 0.65 × 28.96 = 18.8
Using Eq. 1-10, PV = nzRT
Calculating n = 200,000,000 / 379.5
Flow rate under operating conditions =
Gas volume V (= flow rate Q) = 200,000,000 × 0.95 × 10.73 × 530 / (379.5 ×1214.7)
Q = 2,347,000 ft3 per day at operating conditions. Q = 27.16 ft3/s.
Pipe cross sectional area = π r2 = 3.14 × 19 × 19 /(4× 144) = 1.97 ft2
Velocity of gas at the inlet = 27.16/1.97 = 13.8 ft/s.

1.3.11: Viscosity of gases
Viscosity of a fluid relates to the resistance to flow of the fluid. Higher the viscosity, more difficult it is to flow. Viscosity is a number that represents the drag forces caused by the attractive forces in adjacent fluid layers. It might be considered as the internal friction between molecules, separate from that between the fluid and the pipe wall.
The viscosity of a gas is very small compared to that of a liquid. For example, a typical crude oil may have a viscosity of 10 centipoise (cp), whereas a sample of natural gas has a viscosity of 0.0019 cp.
Viscosity may be referred to as absolute or dynamic viscosity measured in cp or kinematic viscosity measured in centistokes (cSt). Other units of viscosity are lb/ft-sec for dynamic viscosity and ft2/s for kinematic viscosity.
Fluid viscosity changes with temperature. Liquid viscosity decreases with increasing temperature, whereas gas viscosity decreases initially with increasing temperature and then increases with further increasing temperature.

Table 1- 6 Viscosity conversion factors

Figure 1-4 can be used to estimate the viscosity of a hydrocarbon gas at various conditions of temperature and pressure if the specific gravity of the gas at standard conditions is known. It is useful when the gas composition is not known. It does not make corrections for H2S, CO2, and N2. It is useful for determining viscosities at high pressure.

1.3.12: The heating value of gases
The heating value of a gas is expressed in Btu/ft3. It represents the quantity of heat in Btu (British Thermal Unit) generated by the complete combustion of one cubic foot of the gas with air at constant pressure at a fixed temperature of 60 0F.
Hydrogen in the fuel burns to water and when the flue gases are cooled to 60°F, the physical state — either vapor or liquid — of this water must be assumed. So the latent heat of vaporization of the water may or may not be considered to be part of the heating value. The result is two definitions for the heating value. The higher or gross heating value, HHV, includes the heat of condensation and the lower or net heating value, LHV, assumes the water remains in the vapor state.
For gas mixture the heating value is calculated as follows:
H = Ʃ xi Hi Eq. 1-20

Example 1-15: Calculate the heating value of gas mixture of Example 1-2

Table 1-7 Solution of Example 1-15

From table 1-7 the Gross calorific value HHV = 1246 Btu/ft3

The higher, ideal, dry heating value of sweet natural gas at 60°F and 760 mm Hg may be calculated with the following equation:
HHV=1568.72 × SG – 2524.88 × XCO2 – 1658.37 × XN2 +141.05 Eq. 1- 21

Applying Eq 1-21 for Example 1-15
The apparent molecular weight= 21.36
Gas Specific gravity = 21.36/28.96 = 0.738
HHV = 1568.72 × 0.738 – 2524.88 × 0.015 – 1658.37 × 0.01 +141.05 = 1244 Btu/ft3

1.4: properties of Hydrocarbon Liquids (Crude Oil)
1.4.1: Introduction
Crude oils are complex mixtures of a vast number of hydrocarbon compounds. Properties of crude petroleum vary appreciably and depend mainly on the origin.
Liquid hydrocarbons are started from Pentane C5 (Natural gasoline) up to solid hydrocarbon (C20) which has a melting point 100 0F. Heavier hydrocarbons (Paraffin and Asphalteen have higher melting points and may be soluble or dispersed in the liquid crude oil depending on solution temperature.)
Crude oil properties depends on its composition which is deferent and variable from crude to another.

Figure 1-4 Hydrocarbon gas viscosity.

1.4.2: Crude oil Density and gravity
Density is defined as the mass of a unit volume of material at a specified temperature. It has the dimensions of grams per cubic centimeter or lb/ft3.
Another general property, which is more widely, is the specific gravity. It is the ratio of the density of oil to the density of water and is dependent on two temperatures, those at which the densities of the oil sample and the water are measured. When the water temperature is 60 0F. The standard temperatures for specific gravity in the petroleum industry is 15/15 0C and 60/60 0F.
Although density and specific gravity are used extensively in the oil industry, the API gravity is considered the preferred property. It is expressed by the following relationship:

0API = 141.5/(Sp.Gr @ 60 Deg F) - 131.5 Eq. 1-22

1.4.3: Crude oil Viscosity.
The best way to determine the viscosity of a crude oil at any temperature is by measurement. If the viscosity is known at only one temperature, Figure 1-5 can be used to determine the viscosity at another temperature by striking a line parallel to that for crudes “A,” “C,” and “D.” Care must be taken to assure that the crude does not have its pour point within the temperature range of interest. If it does, its temperature-viscosity relationship may be as shown for crude “B.”
Solid phase high-molecular-weight hydrocarbons, otherwise known as paraffins, can dramatically affect the viscosity of the crude sample. The cloud point is the temperature at which paraffins first become visible in a crude sample. The effect of the cloud point on the temperature viscosity curve is shown for crude “B” in Figure 1-5. This change in the temperature-viscosity relationship can lead to significant errors in estimation. Therefore, care should be taken when one estimates viscosities near the cloud point.
The pour point is defined as the lowest temperature (5 0F) at which the oil will flow.
The lower the pour point, the lower the paraffin content of the oil.

Figure 1-5, typical viscosity-temperature curves for crude oils. (Courtesy of ASTM D-341.)
(Light crude oil (300–400API), Intermediate crude oil (200–300), & Heavy crude oil (less than 200 API)

In the absence of any laboratory data, correlations exist that relate viscosity and temperature, given the oil gravity. The following equation relating viscosity, gravity, and temperature was developed by Beggs and Robinson after observing 460 oil systems:

µ = 10x -1 Eq. 1-23
where
µ = oil viscosity, cp,
T = oil temperature, 0F,
x = y (T)−1.163,
y = 10z
z = 3.0324 – 0.02023G,
G= oil gravity, API@ 60 0F.
Figure 1-6 is a graphical representation of another correlation.

1.4.4: Oil-Water Mixture Viscosity
The viscosity of produced water depends on the amount of dissolved solids in the water as well as the temperature, but for most practical situations, it varies from 1.5 to 2 centipoise at 500F, 0.7 to 1 centipoise at 1000F, and 0.4 to 0.6 centipoise at 1500F.
When an emulsion of oil and water is formed, the viscosity of the mixture may be substantially higher than either the viscosity of the oil or that of the water taken by themselves. The modified Vand’s equation allows one to determine the effective viscosity of an oil-water mixture and is written in the form
µeff = (1+2.5 ϕ +10 ϕ2) µc Eq. 1- 24
where
µeff = effective viscosity, cp
µc = viscosity of the continuous phase (Oil), cp
Φ = volume fraction of the discontinuous phase (Water).

Figure 1-6, Oil viscosity vs. gravity and temp. (Courtesy of Paragon Eng. Services, Inc.)

1.5: Phase Behavior
1.5.1: Introduction
Before studying the separation of gases and liquids, we need to understand the relationship between the phases. Phase defines any homogeneous and physically distinct part of a system that is separated from other parts of the system by definite bounding surfaces:
The matter has three phases, the simplest example is water.
• Solid (ice),
• Liquid (liquid water),
• Vapor (water vapor).
Solids have a definite shape and are hard to the touch. They are composed of molecules with very low energy that stay in one place even though they vibrate. Liquids have a definite volume but no definite shape. Liquids assume the shape of the container but will not necessarily fill that container. Liquid molecules possess more energy than a solid (allows movement from place to another). By virtue of the energy, there is more space between molecules, and liquids are less dense than solids. Vapors do not have a definite volume or shape and will fill a container in which they are placed. Vapor molecules possess more energy than liquids (very active) and are less dense than liquids.
Our primary concern in this section is the difference in energy level between phases.
Energy is added to melt a solid to form a liquid. Additional energy will cause the liquid to vaporize. One needs to know the phase or phases that exist at given conditions of pressure, volume, and temperature so as to determine the corresponding energy level, to do this we need to study the phase diagram or phase behavior, but first we have to separate components into three classifications:
• Pure substance (single-component systems),
• Two substances,
• Multicomponent.
Phase diagrams illustrate the phase that a particular substance will take under specified conditions of pressure, temperature, and volume.

1.5.2 System Components
Natural gas systems are composed primarily of the lighter alkane series of hydrocarbons, with methane (CH4) and ethane (C2H6) comprising 80% to 90% of the volume of a typical mixture. Methane and ethane exist as gases at atmospheric conditions.
Propane (C3H8), butane (n-C4H10 and i-C4H10), and heavier hydrocarbons may be extracted from the gas system and liquefied for transportation and storage. These are the primary components of liquefied petroleum gas, or LPG.
The intermediate-weight hydrocarbons (pentane through decane) exist as volatile liquids at atmospheric conditions. These components are commonly referred to as pentanes-plus, condensate, natural gasoline, and natural gas liquids (NGL).
Natural gas systems can also contain non-hydrocarbon constituents, including hydrogen sulfide (H2S), carbon dioxide (CO2), nitrogen (N2), and water vapor. These constituents may occur naturally in gas reservoirs, or they may enter the system as contaminants during production, processing, and transportation. In addition, operators may intentionally add odorants, tracers (such as helium), or other components.
Dry, or lean, natural gas systems have high concentrations of the lighter hydrocarbons (methane and ethane), while wet, or rich, gas systems have higher concentrations of the intermediate-weight hydrocarbons. Lean gases burn with a low air-to-gas ratio and display a colorless to blue or yellow flame, whereas rich gases require comparatively higher amounts of air for combustion and burn with an orange flame. Intermediate-weight hydrocarbons may condense from rich gases upon cooling.

Table 1-8 shows typical compositions for a lean gas and a rich gas.

Table 1-8 typical composition of Lean and Rich gases.

1.5.3: Single-Component Systems
A pure component of a natural gas system exhibits a characteristic phase behavior, as shown in Fig. 1-7. Depending on the component’s pressure and temperature, it may exist as a vapor, a liquid, or some equilibrium combination of vapor and liquid

Figure 1-7 P-T Diagram for pure component

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Lines HD, HC, and FH are the equilibrium lines - combinations of pressure and temperature at which the adjoining phases are in equilibrium. At equilibrium, one can change phase, by simply adding or removing energy from the system. Point H, the triple point, is the only combination of pressure and temperature at which all three phases can exist together.
Along line FH no liquid phase is ever present and solid sublimes to vapor. The use of "dry ice" for cooling is an example of this. Line HD is the equilibrium line between solid and liquid. Ice water at 0°C [32°F] and atmospheric pressure occurs on this line. Line HD can have a positive or negative slope depending on whether the liquid expands or contracts on freezing. The energy change occurring along line HD is called the heat of fusion. At any P and T along this line the system can be all solid, all liquid or a mixture of the two depending on the energy level.
This line could be called the solid-liquid saturation or solid-liquid equilibrium line.

Line HC is the saturation or equilibrium curve between vapor and liquid. It starts at the triple point and terminates at the critical point "C." The pressure and temperature conditions at this latter point are known as critical temperature (Tc) and critical pressure (Pc).
At this point the properties of the liquid and vapor phases become identical. For a pure substance the critical point can be defined as that point above which liquid cannot exist as a unique separate phase. Above (Pc), and (Tc), the system is often times referred to as a dense fluid to distinguish it from normal vapor and liquid.
Line HC is often referred to as the vapor pressure curve. Such vapor pressure curves are available from many sources. Line HC is also the bubble point and dew point curve for the pure substance.
The vapor pressure line in Figure 1-8 divides the liquid region from the vapor region.

In Figure 1-7, consider a process starting at pressure P1, and proceeding at constant pressure.
From "m" to "n" the system is entirely solid. The system is all liquid for the segment o-b. At "b" the system is a saturated liquid - any further addition of energy will cause vaporization. At "d," the system is in the saturated vapor state. At temperatures above "d," it is a superheated vapor.
Line HC is thus known by many names - equilibrium, saturated, bubble point, dew point and vapor pressure. For a pure substance these words all mean the same thing.

At the pressure and temperature represented by HC the system may be all saturated liquid, all saturated vapor or a mixture of vapor and liquid.
The rectangle "bfghd" illustrates another important phase property that is confirmed experimentally.
Suppose we place a liquid in a windowed cell at condition "b" and light it so it is easily visible.
We then increase pressure at constant temperature (isothermally). As we proceed toward point “f” the color will begin to fade. At some point, the color disappears completely. The cell now contains what looks like a vapor, but no bubble of vapor was ever seen to form.
At “ f ” (above the critical) the system is in a fourth phase that cannot be described by the senses. It is usually called dense phase fluid, or simply fluid. The word "fluid" refers to anything that will flow and applies equally well to gas and liquid.
This fluid at "f' looks like a gas but possesses different properties from regular gas found to the right of line HC and below the critical pressure. It is denser than regular gas but is more compressible than a regular liquid. “Properties of the liquid and vapor phases become identical”.
Table 1-1 lists Critical pressures and critical temperatures, along with molecular weights, of some pure components present in many natural gas systems.
Figure 1- 8 shows vapor pressure line for light hydrocarbons, where the left part of any component line, represents its liquid phase while the right part represents its gas phase.

Figure 1-8 Vapor pressure for light hydrocarbons.

1.5.4: Multicomponent Systems
In reality, natural gas systems are not pure substances. Rather, they are mixtures of various components, with phase behavior characteristics that differ from those of a single-component system. Instead of having a vapor pressure curve, a mixture exhibits a phase envelope, as shown in Figure 1-9.

Figure 1-9 typical phase envelop of hydrocarbon mixture.

The phase envelope (curve BCD in Figure 1-9) separates the liquid and gas phases. The area within this envelope is called the two-phase region and represents the pressure and temperature ranges at which liquid and gas exist in equilibrium.
The upper line of the two-phase region (curve BC) is the bubble-point line. This line indicates where the first bubble of vapor appears when the pressure of the liquid phase mixture is lowered at constant temperature, or when the temperature increases at constant pressure.
The lower section of the phase envelope (curve CD) is the dewpoint line. When the pressure of a mixture in the gaseous phase is decreased at constant temperature, or when the temperature is lowered at constant pressure, the first drop of liquid forms on this line. The bubble-point line and the dewpoint line meet at the critical point (C).
The highest pressure in the two-phase region is called the cricondenbar, while the highest temperature in the two-phase region is called the cricondentherm.

Figure 1-10, is another example of phase envelope, where:
Cricondenbar - maximum pressure at which liquid and vapor may exist (Point N).
Cricondentherm - maximum temperature at which liquid and vapor may coexist in equilibrium (Point M).
Retrograde Region - that area inside phase envelope where condensation of liquid occurs by lowering pressure or increasing temperature (opposite of normal behavior).
Quality Lines - those lines showing constant percentages which intersect at the critical point (C) and are essentially parallel to the bubble point and dew point curves. The bubble point curve represents 0% vapor and the dew point curve 100% vapor.
Line ABDE represents a typical isothermal retrograde condensation process occurring in a condensate reservoir. Point A represents the single phase fluid outside the phase envelope. As pressure is lowered, Point B is reached where condensation begins. As pressure is lowered further, more liquid forms because of the change in the slope of the quality lines. As the process continues outside the retrograde area, less and less liquid forms until the dewpoint is reached (Point E). Below E no liquid forms.

Figure 1-10 shows another phase envelope for hydrocarbon mixture.

1.5.5: Prediction of phase envelope
The location of the bubblepoint and dewpoint lines may be calculated using vapor-liquid equilibrium (VLE) methods. For most naturally occurring systems above about [2000 psia], the validity of the standard calculation becomes questionable.
The application of K-values to calculate phase quantities and compositions proceeds as follows.
For any stream (F) with mole fractions of components (Z1+Z2+Z3,.., etc.) entering a vessel at certain pressure and temperature, the stream will be divided into Vapor stream(V) with mole fractions of components (Y1+Y2+Y3,.., etc.), and into a liquid phase (L) with mole fractions of components (X1+X2+X3,.., etc.).
Component balance:
Fzi = Vyi + Lxi Eq. 1-25

Figure 1-11 flash separation for hydrocarbon mixture.
where
zi = mol fraction of any component in total feed stream to separation vessel
yi = mol fraction of any component in the vapor phase
xi = mol fraction of any component in the liquid phase
Ki = equilibrium vaporization ratio (equilibrium constant) = yi/xi
F = total mols of feed
V = total mols of vapor
L = total mols of liquid

If we set F = 1.0 so that L and V are now liquid and vapor-to-feed ratios
then zi = Vyi + Lxi
Since yi = Kixi
So, zi = V Ki xi + L xi
xi = zi / ( L + V Ki) Eq. 1-26
Since the summation of liquid fractions must equal one, we can write the following equation.

∑ xi = ∑ zi / ( L + V Ki) = 1 Eq. 1-27

The equation serves as the objective function in an interactive calculation to determine the quantity of L or V. The calculation procedure is as follows:
1. Determine K values of each component at the temperature and pressure of the system.
2. Assume a value of L (remember, V = 1 - L)
3. Solve the equation Eq. 1-27. If ∑xi ≠ 1.0 assume a new value of L and repeat step 2.
4. When ∑ xi = 1.00, the phase quantities L and V are known as well as the liquid phase composition. Vapor phase compositions may be calculated by remembering that yi = Kixi,

The foregoing calculations is known as a flash calculation and is used to predict the equilibrium quantities and compositions of two phase systems.

Special cases of a flash calculation include bubble point (V = 0, L = 1) and dew point (V = 1, L = 0), calculations. Equations for bubblepoint, and dewpoint are as follows:

Bubblepoint condition:

∑ Ki xi = 1.0 Eq. 1-28

Dewpoint condition:

∑ yi/Ki = 1.0 Eq. 1-29

Flash calculation are usually made by computer software, but knowing the basic of calculations is important in understanding the gas-liquid separation process.

Example 1- 16: Calculate the bubblepoint and dewpoint temperature at 250 psia of the following hydrocarbon mixture. Then calculate the amount of vapor and liquid and the composition of the two phases if these feed entered a vessel @ 250 psia and 150 0F.

Table 1-9 hydrocarbon component for example 1-16.

Solution:
Bubblepoint calculation : To calculate the bubblepoint temperature at certain pressure, (All the components are in liquid phase xi = 1).
From eq. 1-28, the bubblepoint will be reached when ∑ Ki xi ≅ 1
Solution Steps:
Assume a temperature value (100 0F), for example.
From the K chart of each compound, find the K value at the system pressure and assumed temperature.
Multiply mole fraction xi of each component by its equilibrium value taken from the table Ki.
Take the sum ∑ Ki xi , if it’s less than 1, choose higher temperature, (1500F for example), and repeat as in the table.
If, ∑ Ki xi is higher than 1, choose a lower temperature.
Repeat till ∑ Ki xi ≅ 1.

Table 1-10 bubblepoint calculation for example 1-16.

We assumed two values of temperature , we found the first value (100 0F) is lower than the bubble point since ∑ Ki xi < 1.00 , and the second value (150 0F) is higher than the bubble point since ∑ Ki xi > 1.00 , the bubble point will be between the two values where ∑ Ki xi ≅ 1.
The Ki values in previous table where collected from “ Design operation and maintenance of gas plants - John Campbell Co.” , since it’s hard to obtain Ki numbers at temperature rather than the pre-drawn temperature lines in K-Charts.
The Values of Ki can be extracted from individual component charts (figures 1-12 to 1-16) (Methane K-chart, Ethane K-chart ….etc.), or can be extracted from “DePriester” chart, fig 1-17.

Dewpoint calculation: To calculate the dewpoint temperature at certain pressure, (All the components are in gas phase yi = 1). From eq. 1-29, the dewpoint will be reached when ∑ yi /Ki ≅ 1
Solution Steps:
Assume a temperature value (150 0F), for example.
From the K chart of each compound, find the K value at the system pressure and assumed temperature.
Divide mole fraction Yi of each component by its equilibrium value taken from the table Ki.
Take the sum ∑ Yi /Ki , if it’s higher than 1, choose higher temperature, (2000F for example), and repeat as in the table.
If, ∑ Yi /Ki is less than 1, choose a lower temperature.
Repeat till ∑ Yi /Ki ≅ 1.

Table 1-11 dewpoint calculation for example 1-16.

We assumed two values of temperature , we found the first value (150 0F) is lower than the dewpoint since ∑ yi /Ki > 1.00 , and the second value (200 0F) is higher than the bubble point since ∑ yi /Ki value is < 1.00 , the dewpoint will be between the two values where ∑ yi /Ki ≅ 1 .
The Ki values in previous table where collected from “ Design operation and maintenance of gas plants - John Campbell Co.” , since it’s hard to obtain Ki numbers at temperature rather than the pre-drawn temperature lines in K-Charts.
The Values of Ki can be extracted from individual component charts (Methane K-chart, Ethane K-chart ….etc.), or can be extracted from “DePriester” chart, fig 1-17.

Flash calculations:
Different values of “L” will be assumed (remember, V = 1 - L), and accordingly Xi will be calculated till we obtain ∑ xi = 1.00.
Ki from chart at 250 psia and 150 0F
Using Eq. 1 -26 xi = zi / ( L + V Ki)

Table 1-12 flash calculations for example 1-16.

The assumed value of L=0.5, found to be lower than the correct value, and the assumed value of L= 0.75 found to be higher than the correct value.
The correct value must be in between the two previous assumed values, and found to be 0.649.
Flash calculations usually performed by computer software, for manual calculations, some K value charts are included in this chapter for the illustration of manual calculations for the previous example. (Figures 1-12 to 1-16)
Other K-values are included in Chapter 25 “Equilibrium Ratio (K) Data” in the “GPSA Engineering Data Book”, or Appendix 5A Volume 1 “Gas conditioning and Processing – The Basic Principles,” Campbell Petroleum Series. In the other hand, the DePriester Chart Figure 1-17, may be used for all hydrocarbon components.

K Value charts:

Figure 1-12 Equilibrium ratio (K) for Methane.

Figure 1-13 Equilibrium ratio (K) for Ethane.

Figure 1-14 Equilibrium ratio (K) for Propane.

Figure 1-15 Equilibrium ratio (K) for i-Butane.

Figure 1-16 Equilibrium ratio (K) for n-Butane.

Figure 1-17 the DePriester (K) Chart for hydrocarbon components.

1.6: Types of Fluid Flow
When a fluid moves through a pipe, two distinct types of flow are possible, laminar and turbulent. Laminar flow occurs in fluids moving with small average velocities and turbulent flow becomes apparent as the velocity is increased above a critical velocity. In laminar flow, the fluid particles move along the length of the pipe in a very orderly fashion, with little or no sideways motion across the width of the pipe.
Turbulent flow is characterized by random, disorganized motion of the particles, from side to side across the pipe as well as along its length. The two types of fluid flow are described by different sets of equations. In general, for most practical situations, the flow will be turbulent.

Figure.1-18. Laminar and turbulent flow in pipes.
1.6.1: Reynolds Number
A useful factor in determining which type of flow is involved is the Reynolds number. This is the ratio of the dynamic forces of mass flow to the shear resistance due to fluid viscosity and is given by the following formula.
Re = VDρ/ µ Eq. 1-30

where
Re = Reynolds number, dimensionless
V = average gas velocity, ft/s
D = pipe inside diameter, ft
ρ = gas density, lb/ft3
μ = gas viscosity, lb/ft-s (1 cp = 0.000672 lb/ft-s)

The flow is considered to be laminar flow when the Reynolds number is below 2000.
Turbulent flow is said to exist when the Reynolds number is greater than 4000. When the Reynolds numbers is between 2000 and 4000, the flow is called critical flow, or undefined flow.
Therefore
Re <= 2000 Flow is laminar
Re > 4000 Flow is turbulent
2000 < Re <= 4000 Flow is critical flow
In terms of the more commonly used units in the gas pipeline industry, the following formula for Reynolds number is more appropriate:
Re = 1.35 x 10-5 (GQ/µd) Eq. 1-31
where
G = gas specific gravity (air = 1)
Q = gas flow rate, standard ft3/day (scfd)
d = pipe inside diameter, in.
μ = gas viscosity, lb/ft-s (1 cp = 0.000672 lb/ft-s)

yasserkassem · Post by **yasserkassem** » Tue Mar 17, 2020 9:32 pm

Two-phase Oil and Gas Separation -Chapter 2 - part 1

Chapter 2

Two-phase Oil and Gas Separation

2.1 Introduction
The production system begins at the wellhead. Fluids produced from oil and gas wells generally constitute mixtures of crude oil, natural gas, and salt water. Crude oil–gas–water mixtures produced from wells, are generally directed, through flow lines and manifold system, to a central processing and treatment facility normally called the gas–oil separation plant (GOSP).
The goal is to attain in the downstream (output) of the “gas oil separation plant”, the following components:
Oil free of water and meets other purchaser’s specifications.
Gas free of hydrocarbon liquid meets other purchaser’s specifications.
Water free of oil and meets environmental, and reservoir regulation for disposal or reinjection.
The first step in processing of the produced stream is the separation of the phases (oil, gas, and water) into separate streams.
Oil may still contain between 10% and 15% water that exists mostly as emulsified water, once initial separation is done, each stream undergoes the proper processing for further field treatment.

2.2 Phase Equilibrium
Equilibrium is a theoretical condition that describes an operating system that has reached a “steady-state” condition whereby the vapor is condensing to a liquid at exactly the same rate at which liquid is boiling to vapor. Simply stated, phase equilibrium is a condition where the liquids and vapors have reached certain pressure and temperature conditions at which they can separate. In most production systems, true equilibrium is never actually reached; however, vapors and liquids move through the system slow enough that a “pseudo” or “quasi” equilibrium is assumed. This assumption simplifies process calculations.
Figure 2-1 illustrates several operating points on a generic phase equilibrium diagram. Point A represents the operating pressure and temperature in the petroleum reservoir. Point B represents the flowing conditions at the bottom of the production tubing of a well. Point C represents the flowing conditions at the wellhead. Typically, these conditions are called flowing tubing pressure (FTP) and flowing tubing temperature (FTT).
Point D represents the surface conditions at the inlet of the first separator.
2.3: Separation process:
The process can be described as:
Two phase separation, or
Three phase separation
The phases referred to are oil, water and gas.
In two phase separation, gas is removed from total liquid (oil plus water).
In three phase separation, however, in addition to the removal of gas from liquids, the oil and water are separated from each other.
Figure 2.2 shows the difference between 2 and 3 phase separation.

2.4: Principles of Physical Separation:
Three principles used to achieve physical separation of gas and liquids or solids are momentum, gravity settling, and coalescing.
Any separator may employ one or more of these principles, but the fluid phases must be "immiscible" and have different densities for separation to occur.

Figure 2-1 Phase equilibrium phase diagram for a typical production system.

2.5: Gravity Separation:
Since a separation depends upon gravity to separate the fluids, the ease with which two fluids can be separated depends upon the difference in the density or weight per unit volume of the fluids. (Density of liquid is much higher than density of gases).
In the process of separating, separation stages are as follows:
1- Separate liquid mist from the gas phase.
2- Separate gas in the form of foam from the liquid phase.
3- In case of 3 phase separation, in addition to the above two requirements, water droplets should be separated from oil phase, and oil droplets should be separated from water phase.
Droplets of liquid mist will settle out from gas, provided:
The gas remains in the separator long enough for mist to drop out.
The flow of the gas through the separator is slow enough that no turbulence occurs, which will keep the gas stream stirred up so that the liquid has no chance to drop out.
The objective of ideal two-phase separation, is to separate the hydrocarbon stream into liquid-free gas and gas-free-liquid. Ideally, the gas and liquid reach a state of equilibrium at the existing conditions of Pressure and Temperature within the vessel.

Figure 2.2 The Difference between 2 & 3 Phase Separation.

Liquid droplets will settle out of a gas phase due to the difference in densities if the gravitational force acting on the droplet is greater than the drag force of the gas flowing around the droplet (see Fig. 2-3). The drag force is the force resulted from the velocity of gas and affecting the entrained droplet of liquid, forcing it to move in the gas flow direction.

Fig. 2-3 A schematic of a force balance on a droplet in a flowing gas stream.
Figures 2-4, and 2-5, illustrates the liquid droplet in gas phase and gas bubble in liquid phase in both configurations of horizontal and vertical separators. From both figures, it’s clear that, in vertical separator, the gravitational settling force is countercurrent or opposite of the drag force resulted from gas movement. While in horizontal separator, the two forces are perpendicular to each other.
The same for the gas bubble entrained in liquid in vertical and horizontal separators.

Fig. 2- 4.The liquid droplet in gas phase and gas bubble in liquid phase in horizontal separator.

Fig. 2- 5.The liquid droplet in gas phase and gas bubble in liquid phase in vertical separator.
2.6: Factors Affecting Separation
Characteristics of the flow stream will greatly affect the design and operation of a separator. The following factors must be determined before separator design:
• Gas and liquid flow rates (minimum, average, and peak),
• Operating and design pressures and temperatures,
• Surging or slugging tendencies of the feed streams,
• Physical properties of the fluids such as density and compressibility factor,
• Designed degree of separation (e.g., removing 100% of mist greater than 10 microns of gas stream),
• Presence of impurities (paraffin, sand, scale, etc.),
• Corrosive tendencies of the liquids or gas.
• Foaming tendencies of the crude oil.
It is important to highlight that: The degree of separation is dependent on the retention time provided. Retention time is affected by the amount of liquid the separator can hold, and the rate at which the fluids enter the vessel.
2.7: Separator categories and nomenclature:
Since, separators is any device of vessel will separate a certain phase from another immiscible phase, there are many types of vessel or devices performing this function, however, their names will differ as follows:
Two- phase separator: A vessel used to separate a mixed-phase stream into gas and liquid phases that are "relatively" free of each other. Other terms used are scrubbers, knockouts, line drips, and decanters.
Flash Tank: A vessel used to separate the gas evolved from liquid flashed from a higher pressure to a lower pressure.
Line Drip: Typically used in pipelines with very high gas-to-liquid ratios to remove only free liquid from a gas stream, and not necessarily all the liquid. Line drips provide a place for free liquids to separate and accumulate.
Liquid-Liquid Separators: Two immiscible liquid phases can be separated using the same principles as for gas and liquid separators. Liquid-liquid separators are fundamentally the same as gas-liquid separators except that they must be designed for much lower velocities. Because the difference in density between two liquids is less than between gas and liquid, separation is more difficult.
Scrubber or Knockout: A vessel designed to handle streams with high gas-to-liquid ratios. The liquid is generally entrained as mist in the gas or is free-flowing along the pipe wall. These vessels usually have a small liquid collection section. The terms are often used interchangeably.
Slug Catcher: A particular separator design able to absorb sustained in-flow of large liquid volumes at irregular intervals.
Usually found on gas gathering systems or other two phase pipeline systems. A slug catcher may be a single large vessel or a manifolded system of pipes.
Three Phase Separator: A vessel used to separate gas and two immiscible liquids of different densities (e.g. gas, water, and oil).
Filter Separators: A filter separator usually has two compartments.
The first compartment contains filter-coalescing elements. As the gas flows through the elements, the liquid particles coalesce into larger droplets and when the droplets reach sufficient size, the gas flow causes them to flow out of the filter elements into the center core. The particles are then carried into the second compartment of the vessel (containing a vane-type or knitted wire mesh mist extractor) where the larger droplets are removed. A lower barrel or boot may be used for surge or storage of the removed liquid.
2.8: Functional Sections of a Gas-Liquid Separator
Regardless of the size or shape of a separator, each gas-liquid separator contains four major sections. Figures 2-7 and 2-8 illustrate the four major sections of a horizontal and vertical two-phase separator, respectively.

Fig. 2- 6.Gas liquid separation selection map.

2.8.1: Inlet Diverter Section
The inlet stream to the separator is typically a high-velocity turbulent mixture of gas and liquid. Due to the high velocity, the fluids enter the separator with a high momentum. Collision or abruptly changes the direction of flow by absorbing the momentum of the liquid and allowing the liquid and gas to separate. This results in the initial “gross” separation of liquid and gas. The inlet diverter, sometimes referred to as the primary separation section. Therefor this section is used to reduce the momentum of the inlet flow stream, perform an initial bulk separate ion of the gas and liquid phases, and enhance gas flow distribution. There are varieties of inlet devices available and these will be discussed in more detail in a later section.
2.8.2: Liquid Collection Section
The liquid collection section, located at the bottom of the vessel, it acts as a receiver for all liquid removed from the gas in the inlet, gas gravity, and mist extraction sections. The liquid collection section provides the required retention time necessary for any entrained gas in the liquid to escape to the gravity settling section. In addition, it provides a surge volume to handle intermittent slugs.
In three-phase separation applications, the liquid gravity section also provides residence time to allow for separation of water droplets from a lighter hydrocarbon liquid phase and vice-versa. Due to the smaller difference in gravity between crude oil and water, compared to gas and liquid in two-phase separation, Liquid-liquid separation requires longer retention times than gas-liquid separation.
Also in in three phase separators, a coalescing packs are sometimes used to promote hydrocarbon liquid – water separation, though they should not be used in applications that are prone to plugging, e.g. wax, sand, etc.

2.8.3: Gravity Settling Section
As the gas stream enters the gravity settling section, its velocity drops and small liquid droplets that were entrained in the gas and not separated by the inlet diverter are separated out by gravity and fall to the gas liquid interface, preconditioning the gas for final polishing by the mist extractor.
. The gravity settling section is sized so that liquid droplets greater than 100 to 140 microns fall to the gas-liquid interface while smaller liquid droplets remain with the gas. Liquid droplets greater than 100 to 140 microns are undesirable as they can overload the mist extractor at the separator outlet.
In some horizontal designs, straightening vanes are used to reduce turbulence. The vanes also act as droplet coalescers, which reduces the horizontal length required for droplet removal from the gas stream.

2.8.4: Mist Extractor Section
Gas leaving the gravity settling section contains small liquid droplets, generally less than 100 to 140 microns. Before the gas leaves the vessel, it passes through a coalescing section or mist extractor. This section uses coalescing elements that provide a large amount of surface area used to coalesce and remove the small droplets of liquid. As the gas flows through the coalescing elements, it must make numerous directional changes. Due to their greater mass, the liquid droplets cannot follow the rapid changes in direction of flow. These droplets impinge and collect on the coalescing elements, where they fall to the liquid collection section. Quoted liquid carryover from the various types of mist extraction devices are usually in the range of 0.1 - 1 gal/MMscf.

Fig. 2- 7.Horizontal Separator sections with gas bubble in liquid phase, and liquid droplet in gas phase.

2.9: Separator Configurations
Factors to be considered for separator configuration selection include:
• What separation quality is required by downstream equipment and processes?
• How well will extraneous material (e.g. sand, mud, corrosion products) be handled?
• How much plot space will be required?
• Will the separator be too tall for transport if skidded?
• Is there enough interface surface for 3-phase separation (e.g. gas/hydrocarbon/glycol liquid)?
• Can heating coils or sand jets be incorporated if required?
• How much surface area is available for degassing of separated liquid?
• Must surges in liquid flow be handled without large changes in level?
• Is large liquid retention volume necessary?
• What are the heat retention requirements (e.g. freeze protection)?

Fig. 2- 8.Vertical Separator sections with gas bubble in liquid phase and liquid droplet in gas phase.

2.10: Types of Separators
Separators are usually characterized by orientation as vertical or horizontal. They may be further classified as two-phase (gas-liquid) or three-phase (gas-liquid-liquid). Horizontal separators can be single- or double-barrel and can be equipped with sumps or boots.
Each configuration has specific advantages and limitations. Selection is based on obtaining the desired results at the lowest “life-cycle” cost.

2.10.1: Vertical Separators
Vertical separators, shown in Fig. 2-9, are usually selected when the gas-liquid ratio is high or total gas volumes are low. In a vertical separator, the fluids enter the vessel through an inlet device whose primary objectives are to achieve efficient bulk separation of liquid from the gas and to improve flow distribution of both phases through the separator. Liquid removed by the inlet device is directed to the bottom of the vessel.
The gas moves upward in the gravity settling section, where the liquid droplets fall vertically downward counter-current to the upward gas flow. The settling velocity of a liquid droplet is directly proportional to its diameter. If the size of a liquid droplet is too small, it will be carried up and out with the vapor.
Thus, a mist extractor section is added to capture small liquid droplets.
Liquid removed by the mist extractor is coalesced into larger droplets that then fall through the gas to the liquid reservoir in the bottom. Liquid continues to flow downward through liquid collection section to the liquid outlet. As the liquid reaches equilibrium, gas bubbles flow counter to the direction of the liquid flow and eventually migrate to the vapor space.

The ability to handle liquid slugs is typically obtained by increasing vessel height to accommodate additional surge volume. Level control is normally not highly critical and liquid level can fluctuate several inches without affecting the separation performance or capacity of the vessel.
Typical vertical separator L/D ratios are normally in the 2–4 range.

Fig. 2-9 Vertical Two-phase separator.

The pressure in the separator is maintained by a pressure controller mounted on the gas outlet. The pressure controller senses changes in the pressure in the separator and sends a signal to either open or close the pressure control valve accordingly. By controlling the rate at which gas leaves the vapor space of the vessel, the pressure in the vessel is maintained.
The liquid dump valve is regulated by a level controller. The level controller senses changes in liquid level and controls the dump valve accordingly.
There are seldom any internals in the liquid collection section except possibly a still well for the level control float or displacer. The still well usually consists of walled box or tube, open on the top and bottom. Its function is to stop wave action in the separator from interfering with the level controller’s operation.
Vertical separators are well suited for production containing sand and other sediment and thus are often fitted with a false cone bottom to handle sand production.
As an example of a vertical separator, consider a compressor suction scrubber. In this service the vertical separator:
• Does not need significant liquid retention volume
• A properly designed liquid level control loop responds quickly to any liquid that enters, thus avoiding tripping an alarm or shutdown
• The separator occupies a small amount of plot space

2.10.2: Horizontal Separators
Horizontal separators are most efficient when large volumes of liquid are involved. They are also generally preferred for three-phase separation applications. In a horizontal separator, shown in Fig. 2-10. The fluid enters the separator and hits an inlet diverter, causing a sudden change in momentum. The initial gross separation of liquid and vapor occurs at the inlet diverter. The force of gravity causes the liquid droplets to fall out of the gas stream to the bottom of the vessel, where it is collected.
The liquid collection section provides the retention time required to let entrained gas evolve out of the oil and rise to the vapor space and reach a state of “equilibrium.” It also provides a surge volume, if necessary, to handle intermittent slugs of liquid. The liquid leaves the vessel through the liquid dump valve.

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Fig. 2-10 Horizontal Two-phase separator.

Gas and oil mist flow over the inlet diverter and then horizontally through the gravity settling section above the liquid. As the gas flows through this section, small droplets of liquid that were entrained in the gas and not separated by the inlet diverter are separated out by gravity and fall to the gas-liquid interface.
Some of the drops are of such a small diameter that they are not easily separated in the gravity settling section. Before the gas leaves the vessel, it passes through a coalescing section or mist extractor. This section uses elements of vanes, wire mesh, or plates to provide a large amount of surface area used to coalesce and remove the very small droplets of liquid in one final separation before the gas leaves the vessel.
Normally, horizontal separators are operated half full of liquid to maximize the surface area of the gas-liquid interface.
Horizontal separators have certain advantages with respect to gravity separation performance in that the liquid droplets or gas bubbles are moving perpendicular to the bulk phase velocity, rather than directly against it as in vertical flow, which makes separation easier.
The level controller and liquid dump valve operate the same as in a vertical separator.
Pressure and level are maintained as in a vertical separator.
Horizontal separators are smaller and thus less expensive than a vertical separator for a given gas and liquid flow rate. Horizontal separators are commonly used in flow streams with high gas-liquid ratios and foaming crude.
Typical L/D ratios for horizontal separators normally fall in the range of 2.5–5.

2.10.3: Double-Barrel Horizontal Separators
A double-barrel horizontal separator is a variation of the horizontal separator Figure 2-11. Double-barrel horizontal separators are commonly used in applications where there are high gas flow rates and where there is a possibility of large liquid slugs, e.g., slug catchers.
Single-barrel horizontal separators can handle large gas flow rates but offer poor liquid surge capabilities. The double-barrel horizontal separator partially alleviates this shortcoming. In these designs the gas and liquid chambers are separated as shown in Figure 2-11. The flow stream enters the vessel in the upper barrel and strikes the inlet diverter. The gas flows through the gravity settling section, where it may encounters a baffle type mist extractors, or directly to the wire mish mist extractor to en route to the gas outlet.

Fig. 2-11 Double-barrel Two-phase separator.

The baffles help the free liquids to fall to the lower barrel through flow pipes. The liquids drain through a flow pipe or equalizing tube into the lower barrel. Small amounts of gas entrained in the liquid are liberated in the liquid collection barrel and flow up through the flow pipes or equalizing tubes. In this manner the liquid accumulation is separated from the gas stream so that there is no chance of high gas velocities re-entraining liquid as it flows over the interface. Because of their additional cost, and the absence of problems with single-vessel separators, they are not widely used in oil field systems. However, in gas handling, conditioning, and processing systems, two-barrel separators are typically used as gas scrubbers on the inlet of compressors, glycol contact towers, and gas treating systems where the liquid flow rate is extremely low relative to the gas flow rate.

2.10.4: Horizontal Separator with a “Boot” or “Water Pot”
A single barrel separator with a liquid “boot” or “water pot” at the outlet end is a special case of a two-barrel separator (Figure 2-12). It is. The main body of the separator operates essentially dry as in a two-barrel separator. The small amounts of liquid in the bottom flow to the boot end, which provides the liquid collection section. These vessels are less expensive than two-barrel separators, but they also contain less liquid handling capability. It is used when there are very low liquid flow rates.
In applications where there is very little liquid flow, often a horizontal separator will be designed with a liquid sump on the outlet end to provide the required liquid retention time. This results in an overall smaller diameter for the vessel.

Fig. 2-12 Horizontal two-phase separator with boot “Water pot”.

2.10.5: Filter Separators
Filter separator is another type of separator that is frequently used in some high-gas/low liquid flow applications. They can be either horizontal or vertical in configuration. Filter separators are designed to remove small liquid and solid particles from the gas stream. These units are used in applications where conventional separators employing gravitational or centrifugal force are ineffective. Figures 2-13A, B, and C. show a horizontal two barrel filter separator design. Filter tubes in the initial separation section cause coalescence of any liquid mist into larger droplets as the gas passes through the tubes. A secondary section of vanes or other mist extractor elements removes these coalesced droplets. Filter separators are commonly used on compressor inlets in field compressor stations, final scrubbers upstream of glycol contact towers, and instrument/fuel gas applications. The design of filter separators is proprietary and dependent upon the type of filter element employed. Some filter elements can remove 100% of 1-micron particles and 99% of 1/2-micron particles when they are operated at rated capacity and recommended filter-change intervals.
Filter element is consists of a perforated metal cylinder with gasketed ends for compression sealing. A fiberglass cylinder, typical 1/2-inch thick, surrounds the perforated metal cylinder. Gas flow is from outside the fiberglass cylinder to the center of the perforated metal tube.

Fig. 2-13A Horizontal double barrel two-phase filter separator.

Fig. 2-13B Horizontal double barrel two-phase filter separator.

Fig. 2-13C Horizontal double barrel two-phase filter separator.
2.10.6: Scrubbers
A scrubber is a two-phase separator (similar to fig. 2-9) that is designed to recover liquids carried over from the gas outlets of production separators or to catch liquids condensed due to cooling or pressure drops. Liquid loading in a scrubber is much lower than that in a separator. Typical applications include: upstream of mechanical equipment such as compressors that could be damaged, destroyed or rendered ineffective by free liquid; downstream of equipment that can cause liquids to condense from a gas stream (such as coolers); upstream of gas dehydration equipment that would lose efficiency, be damaged, or be destroyed if contaminated with liquid hydrocarbons; and upstream of a vent or flare outlet.
Vertical scrubbers are most commonly used. Horizontal scrubbers can be used, but space limitations usually dictate the use of a vertical configuration.

2.10.7: Slug Catchers
A slug catcher, is a special case of two-phase gas-liquid separator that is designed to handle large gas capacities and liquid slugs on a regular basis, and it is commonly used in gas gathering pipelines. Liquid are usually accumulate in the bottom of gas pipes, especially in offshore pipes, accumulation will continue till the pressure upstream it will be higher enough to push all the liquid as a slug to the gathering station. Such slugs will disturb any production separator in process area, so a slug catcher is necessary to handle the gas and liquid slugs from time to time. Also slug catcher is necessary to handle slugs resulting from pip pigging which is periodically used to sweep the lines of liquids.
When the pigs sweep the liquids out of the gathering lines, large volumes of liquids must be handled by the downstream separation equipment.
There are numerous slug catcher designs. Figures 2-14, and 2-15 are both two-phase horizontal slug catcher with liquid “fingers.” Gas and liquid slug from the gathering system enters the horizontal portion of the two-phase vessel, where primary gas-liquid separation is accomplished. Gas exits the top of the separator through the mist extractor while the liquid exits the bottom of the vessel through a series of large-diameter tubes or “fingers.” The tubes provide a large liquid holding volume and routes the liquid to a three-phase free-water knockout (FWKO) for further liquid-liquid separation. The design of an FWKO is discussed in next chapter.

Fig. 2-14. Two-phase horizontal slug catcher with liquid “fingers.”

Fig. 2-15. Two-phase horizontal slug catcher.

2.11: Selection Considerations
The geometry of and physical and operating characteristics give each separator type advantages and disadvantages.
Horizontal separators have the following features they are:
Smaller,
More efficient at handling large volumes of gas, and less expensive than vertical separators for a given gas capacity.
In the gravity settling section of a horizontal vessel, the liquid droplets fall perpendicular to the gas flow and thus are more easily settled out of the gas continuous phase. Also, since the interface area is larger in a horizontal separator than a vertical separator, it is easier for the gas bubbles, which come out of solution as the liquid approaches equilibrium, to reach the vapor space.
Offer greater liquid capacity and are best suited for liquid-liquid separation and foaming crude. Thus, from a pure gas/liquid separation process, horizontal separators would be preferred.

However, they do have the following drawbacks, which could lead to a preference for a vertical separator in certain situations:
Horizontal separators are not as good as vertical separators in handling solids. The liquid dump line of a vertical separator can be placed at the center of the bottom head so that solids will not build up in the separator but continue to the next vessel in the process. As an alternative, a drain could be placed at this location so that solids could be disposed of periodically while liquid leaves the vessel at a slightly higher elevation.
In a horizontal vessel, it is necessary to place several drains along the length of the vessel. Since the solids will have an angle of repose of 450 to 600, the drains must be spaced at very close intervals. Attempts to lengthen the distance between drains, by providing sand jets in the vicinity of each drain to fluidize the solids while the drains are in operation, are expensive and have been only marginally successful in field operations.
Horizontal vessels require more plan area to perform the same separation as vertical vessels. While this may not be of importance at a land location, it could be very important offshore.
The ability of a separator to absorb a slug of liquid is called the surge capacity of a separator. Horizontal vessels can have less liquid surge capacity than vertical vessels sized for the same steady-state flow rate. For a given change in liquid surface elevation, there is typically a larger increase in liquid volume for a horizontal separator than for a vertical separator sized for the same flow rate. However, the geometry of a horizontal vessel causes any high level shut-down device to be located close to the normal operating level. In very large diameter [greater than 6 ft.] horizontal vessels and in vertical vessels, the shut-down device could be placed much higher, allowing the level controller and dump valve more time to react to the surge. In addition, surges in horizontal vessels could create internal waves, which could activate a high-level sensor prematurely.

It should be pointed out that vertical vessels also have some drawbacks that are not process related and must be considered in making a selection.
These are as follows:
The relief valve and some of the controls may be difficult to service without special ladders and platforms.
The vessel may have to be removed from a skid for trucking due to height restrictions.
Generally, horizontal separators are less expensive than equally sized vertical separators. Since vertical separators are supported only by the bottom skirt (refer to Figure 2-16), the walls of vertical separators must be somewhat thicker than a similarly sized and rated horizontal separator, which may be supported by saddles. Furthermore, large vertical separators, when exposed to high winds, can be subjected to large lateral (wind) loads. When this is the case, the vertical separator’s wall thickness must be increased, which in turn increases the cost of the overall vessel.

Overall, horizontal vessels are the most economical for normal oil-gas separation, particularly where there may be problems with emulsions, foam, or high gas-oil ratios (GOR). Vertical vessels work most effectively in low-GOR applications.
They are also used in some very high GOR applications, such as scrubbers where only fluid mists are being removed from the gas and where extra surge capacity is needed to allow shutdown to activate before liquid is carried out the gas outlet (e.g., compressor suction scrubber).

Fig. 2- 16. Comparison of vertical and horizontal support structures.
2.12: Internal Vessel Components
2.12.1: Inlet Diverters
Inlet diverters serve to impart flow direction of the entering vapor/liquid stream and provide primary separator between the liquid and vapor. There are many types of inlet diverters as shown in fig. 2-17.
• No inlet device
• Diverter plate
• Half-pipe
• Reversed pipe (elbow)
• Dished head
• Vane-type
• Cyclonic

Figures 2-18, 2-19, and 2-20 show baffle plates, vane, and centrifugal diverters.
The main functions of the inlet device are:
• Reduce the momentum of the inlet stream and enhance flow distribution of the gas and liquid phases.
• Efficient separation of the bulk liquid phase.
• Prevent droplet shattering and re-entrainment of bulk liquid phase.
There are several different types of separator inlet devices that are commonly used:

Fig. 2-17. Main types of inlet diverters.

Fig. 2-18. Baffle types inlet diverters.

A baffle plate can be a spherical dish, flat plate, angle iron, cone, elbow, or just about anything that will accomplish a rapid change in direction and velocity of the fluids and thus disengage the gas and liquid.
At the same velocity the higher-density liquid possesses more energy and, thus, does not change direction or velocity as easily as the gas.
Thus, the gas tends to flow around the diverter while the liquid strikes the diverter and then falls to the bottom of the vessel. The design of the baffles is governed principally by the structural supports required to resist the impact-momentum load. The advantage of using devices such as a half-sphere elbow or cone is that they create less disturbance than plates or angle iron, cutting down on re-entrainment or emulsifying problems.
Centrifugal inlet diverters use centrifugal force, rather than mechanical agitation, to disengage the oil and gas. These devices can have a cyclonic chimney or may use a tangential fluid race around the walls (refer to Figure 2-21). Centrifugal inlet diverters are generally use an inlet nozzle sufficient to create a fluid velocity of about 20 ft/s around a chimney. Centrifugal diverters can be designed to efficiently separate the liquid while minimizing the possibility of foaming or emulsifying problems. The disadvantage is that their design is rate sensitive. At low velocities they will not work properly. Thus, they are not normally recommended for producing operations where rates are not expected to be steady.
In addition to the inlet device itself, it has been determined that the inlet piping configuration is also important. The vane type and cyclonic inlet devices generally provide improved separation performance compared to the others.
A Comparison of different inlet diverters efficiency, are in table 2-1.

Fig. 2-19. Vane type inlet diverters.

Fig. 2-20. Cyclone type inlet diverters.

Table 2-1. Inlet diverters comparison.

Figure 2-21. Centrifugal inlet diverters. (Top) Cyclone baffle. (Bottom) Tangential raceway.

2.12.2: Wave Breakers
In long horizontal vessels, waves may result from surges of liquids entering the vessel or will result if the horizontal vessel is located on a floating structure. Wave breakers are nothing more than perforated baffles or plates that are placed perpendicular to the flow located in the liquid collection section of the separator. These baffles dampen any wave action that may be caused by incoming fluids. The wave actions in the vessel must be eliminated so level controls, level switches, and weirs may perform properly. Figure 2-22 is a three-dimensional view of a horizontal separator fitted with an inlet diverter, de-foaming element, mist extractor, and wave breakers.

Figure 2-22. A separator fitted with an inlet diverter, defoaming element, mist extractor, & wave breaker

2.12.3: Defoaming Plates
Foam at the interface may occur when gas bubbles are liberated from the liquid. Foam can severely degrade the performance of a separator. This foam can be destabilized with the addition of chemicals at the inlet, but the more effective solution is to force the foam to pass through a series of inclined parallel plates or tubes as shown in Figure 2-23.
These closely spaced, parallel plates or tubes provide additional surface area, which breaks up the foam and allows the foam to collapse into the liquid layer.

Figure 2-23. defoaming element in horizontal separator.

2.12.4: Vortex Breaker
Liquid leaving a separator may form vortices or whirlpools, which can pull gas down into the liquid outlet. Therefore, horizontal separators are often equipped with vortex breakers, which prevent a vortex from developing when the liquid control valve is open. A vortex could suck some gas out of the vapor space and re-entrain it in the liquid outlet( refer to fig. 2-24A. One type of vortex breaker is shown in Figure 2-24B. It is a covered cylinder with radially directed flat plates. As liquid enters the bottom of the vortex breaker, any circular motion is prevented by the flat plates. Any tendency to form vortices is removed. Figure 2-25 illustrates other commonly used vortex breakers.

Figure 2-24A. Vortexing of liquids

Figure 2-24B. Vortex breaker.
2.12.5: Stilling Well
A stilling well, which is simply a slotted pipe fitting surrounding an internal level control displacer, protects the displacer from currents, waves, and other disturbances that could cause the displacer to sense an incorrect level measurement.

Figure 2-25. Vortex breakers.
2.12.6: Sand Jets and Drains
In horizontal separators, one worry is the accumulation of sand and solids at the bottom of the vessel. If allowed to build up, these solids will upset the separator operations by taking up vessel volume. Generally, the solids settle to the bottom and become well packed.
To remove the solids, sand drains are opened in a controlled manner, and then high-pressure fluid, usually produced water, is pumped through the jets to agitate the solids and flush them down the drains. The sand jets are normally designed with a 20-ft/s, jet tip velocity and aimed in such a manner to give good coverage of the vessel bottom.

Figure 2-26. Schematic of a horizontal separator fitted with sand jets and inverted trough.

Figure 2-27. Schematic of a horizontal separator fitted with sand jets and inverted trough.

2.12.7: Mist Extractors
2.12.7.1: Introduction
Mist extractors or mist eliminators or demister, are names of an equipment used to remove the liquid droplets and solid particles from the gas stream.
All mist extractor types are based on the some kind of intervention in the natural balance between gravitational and drag forces. This is accomplished in one or more of the following ways:
• Overcoming drag force by reducing the gas velocity (gravity separators or settling chambers)
• Introducing additional forces (venturi scrubbers, cyclones.)
• Increasing gravitational force by boosting the droplet size (impingement-type)
The following factors should be considered before selection:
• Size of droplets the equipment must remove
• Accepted pressure drop across the mist extractor
• Susceptibility of the equipment to plugging by solids, if solids are present
• Liquid handling capability of the equipment
• Whether the mist extractor/eliminator can be installed inside existing vessel, or if it requires a standalone vessel instead
• Cost of the mist extractor/eliminator itself and required vessels, piping, instrumentation, and utilities

2.12.7.2: Impingement-Type Mist Extractor
Impingement-type mist extractor is the most widely used type of mist extractors because it offers good balance between efficiency, operating range, pressure drop requirement, and installed cost. These types consist of baffles, wire meshes, and micro-fiber pads. Impingement-type mist extractors may involve just a single baffle or disc installed in a vessel. As illustrated in Figure 2-28, as the gas approaches the surface of the baffle or disc (commonly referred to as a target), fluid streamlines spread around the baffle or disc. The higher the stream velocity, the closer to the target these streamlines start to form. A droplet can be captured by the target in an impingement-type mist extractor/eliminator via any of the following three mechanisms: inertial impaction, direct interception, and diffusion (Fig. 2-28A and B).

Figure 2-28A. The three primary mechanisms of mist capture via impingement are inertial impaction, direct interception, and Brownian diffusion.

Figure 2-28B. The three primary mechanisms of mist capture via impingement are inertial impaction (left), direct interception (center), and Brownian diffusion (right).

• Inertial impaction. Because of their mass, particles 1 to 10 microns in diameter in the gas stream have sufficient momentum to break through the gas streamlines and continue to move in a straight line until they impinge on the target. Impaction is generally the most important mechanism in wire mesh pads and impingement plates.
• Direct interception. There are also particles in the gas stream that are smaller, between 0.3 to 1 microns in diameter, than those above.
These do not have sufficient momentum to break through the gas streamlines. Instead, they are carried around the target by the gas stream. However, if the streamline in which the particle is traveling happens to lie close enough to the target so that the distance from the particle centerline to the target is less than one-half the particle’s diameter, the particle can touch the target and be collected. Interception effectiveness is a function of pore structure. The smaller the pores, the greater the media to intercept particles.
• Diffusion. Even smaller particles, usually smaller than 0.3 microns in diameter, exhibit random Brownian motion caused by collisions with the gas molecules. This random motion will cause these small particles to strike the target and be collected, even if the gas velocity is zero. Diffusion is favored by low velocity and high-concentration gradients.

2.12.7.3: Baffles (Vane Type) mist extractor
This type of impingement mist extractor consists of a series of baffles, vanes, or plates between which the gas must flow. The most common is the vane or chevron-shape, as shown in Figures 2-29, 2-30, and 2-31. The vanes force the gas flow to be laminar between parallel plates that contain directional changes. The surface of the plates serves as a target for droplet impingement and collection. A number of different vane pack designs are available. Pack thicknesses are generally in the range of 6–12 inches. Vanes are usually arranged in a zig-zag or sinusoidal pattern, The space between the baffles ranges from 5 to 75 mm, with a total depth in the flow direction of 150 to 300 mm.
Figures 2-32 and 2-33 illustrate a vane mist extractor installed in a vertical and horizontal separator, respectively. Figure 2-34 shows a vane mist extractor made from an angle iron. Figure 2-35 illustrates an “arch” plate mist extractor. As gas flows through the plates, droplets impinge on the plate surface. The droplets coalesce, fall, and are routed to the liquid collection section of the vessel. Vane-type eliminators are sized by their manufacturers to assure both laminar flow and a certain minimum pressure drop. Vane or chevron-shaped mist extractors remove liquid droplets 10 to 40 microns and larger.

Figure 2-29. Vane-type mist extractor.

Figure 2-30. Vane-type element with corrugated plates and liquid drainage trays.

Figure 2-31. Vane-type mist extractor/eliminator.

Figure 2-32. Typical vane-type mist extractor installed in vertical separator.

Figure 2-33. Vane-type mist extractor installed in horizontal separator.

Figure 2-34. A vane-type mist extractor made from angle iron.

Figure 2-35. An “arch” plate-type mist extractor.

Separation Performance—
The operation and performance is usually dictated by a design velocity expressed as follows:

Vt = K [(ρl - ρg ) / ρg]0.5 Eq. 2-1

where
V = gas velocity, ft/s
K = Souders–Brown coefficient,
ρl = liquid or droplet density, lb/ft3
ρg = gas density, lb/ft3

The “K” factor or Souders–Brown coefficient, is determined experimentally for each plate geometry. Its value ranges from 0.3 to 1.0 ft/s in typical designs. Since impaction is the primary collection mechanism, at too low a value of “K” the droplets can remain in the gas streamlines and pass through the device uncollected. The upper limit is set to minimize re-entrainment, which is caused either by excessive breakup of the droplets as they impinge onto the plates or by shearing of the liquid film on the plates.
Table. 2-2 provides a summary of performance parameters.

Table. 2-2. typical vane pack separation performance.

Higher gas velocities can be handled if the vanes are installed in a horizontal gas flow, instead of vertical up-flow. In the horizontal configuration the liquid can easily drain downward due to gravity and thus out of the path of the incoming gas, which minimizes re-entrainment of the liquid.
Recently developed hollow vane designs with interconnected liquid drainage passages are capable of high gas handling capacities in a vertical upflow orientation.
The vane type appears most often in process systems where the liquid entrainment is contaminated with solids, or where high liquid loading exists. Vane-type mist extractors are less efficient in removing very small droplets than other impaction-types such as wire mesh or micro-fiber.
Standard designs are generally limited to droplets larger than 40 microns.
However, high-efficiency designs provide droplet removal down to less than 15 microns in diameter. Vane packs typically have pressure drops in the range of 0.5–3.5 inches of water.

Vane packs show a drop-off of removal efficiency as pressure increases. This is primarily a result of the decreasing allowable gas velocity with increasing pressure caused mainly by increased gas density.

Mesh pads also rely on velocity/droplet inertia to remove liquid droplets via impingement but they are less susceptible to capture efficiency reduction than vane packs because mesh pads have far more collection “targets”, i.e. wire/fiber filaments.
Turndown is generally more of an issue with vane-packs, with droplet removal efficiency decreasing measurably as velocity decreases from design.
Vane-type mist extractors are also impacted by inlet liquid loading, but generally have considerably more tolerance towards liquids than mesh-pads.

The required mist extractor area is obtained from

A = Qg / Vt Eq. 2-2

where
A = area of mist extractor (ft2)
Qg = actual gas flow rate, ft3/sec

2.12.7.4: Wire-Mesh mist extractor
Wire-mesh mist extractors, or pads, are made by knitting wire, metal or plastic, into tightly packed layers which are then crimped and stacked to achieve the required pad thickness. If removal of very small droplets, i.e. less than 10 micron, is required, much finer fibers may be interwoven with the primary mesh to produce a co-knit pad. Mesh pads remove liquid droplets mainly by impingement of droplets onto the wires and/or co-knit fibers followed by coalescence into droplets large enough to disengage from the bottom of the pad and drop through the rising gas flow into the liquid holding part of the separator. Mesh pads are not recommended for dirty or fouling service as they tend to plug easily.
Wire-mesh is the most common type of mist extractor found in production operations (Figure 2-36).
Most installations will use a 6-inch thick pad with 9-12 lb/ft3 bulk density. Minimum recommended pad thickness is 4 inches. They are usually constructed from wires of diameter ranging from 0.10 to 0.28 mm, with a typical void volume fraction of 0.95 to 0.99. The wire pad is placed between top and bottom support grids near the outlet of a separator, generally on a support ring (vertical separator) or frame (horizontal separator). (Figures 2-37 and 2-38.)
Wire-mesh mist extractors are normally installed in vertical upward gas flow, although horizontal flows are employed in some specialized applications. In a horizontal flow the designer must be careful because liquid droplets captured in the higher elevation of the vertical mesh may drain downward at an angle as they are pushed through the mesh, resulting in re-entrainment.

Figure 2-36. Example wire-mesh mist extractor.

Whether installed inside a piece of process equipment or placed inside a separate vessel of its own, a wire-mesh or baffle-type mist extractor offers low-pressure drop. To ensure a unit’s operation at design capacity and high mist elimination efficiency, the flow pattern of the gas phase must be uniform throughout the element.
When there are size limitations inside a process vessel, an integral baffle plate can be used on the downstream side face of the wire-mesh element as a vapor distributor. When knockout drums are equipped with vanes or wire-mesh pads, one can use any one of the four following design configurations: horizontal or vertical vessels, with horizontal or vertical vane or mesh elements.
The classic configuration is the vertical vessel with horizontal element.
In order to achieve uniform flow, one has to follow a few design criteria (refer to Figure 2-39).
The effectiveness of wire-mesh depends largely on the gas being in the proper velocity range [refer to Eq. (2-1)]. If the velocities are too high, the liquids knocked out will be re-entrained. If the velocities are low, the vapor just drifts through the mesh element without the droplets impinging and coalescing. The lower limit of the velocity is normally set at 30% of design velocity, which maintains a reasonable efficiency. The upper limit is governed by the need to prevent re-entrainment of liquid droplets from the downstream face of the wire-mesh device. A properly sized wire-mesh unit can remove 100% of liquid droplets larger than 3 to 10 microns in diameter.

Figure 2-37. Vertical separators fitted with wire-mesh pads supported by support rings.

Figure 2-38. Horizontal separator fitted with wire-mesh pads supported by a frame.

Separation Performance — There are two main aspects to mesh pad separation performance.
• droplet removal efficiency
• gas handling capacity
Droplet removal efficiency is typically given by the manufacturer as a curve showing % removal as a function of droplet size at design flow and a nominal liquid loading. These curves are usually based on tests of an air-water system at atmospheric pressure.
The gas capacity of mesh pads is almost universally specified by a load or sizing factor, K, as utilized in the Souders and Brown equation given by Eq. 2-1:
Vt = K [(ρl - ρg ) / ρg]0.5 Eq. 2-1
The required mist extractor area is obtained from Eq. 2-2.
A = Qg / Vt Eq. 2-2

The design K value provides a certain degree of margin before liquid entrainment/carryover becomes excessive. Efficiency and capacity are normally inversely related, i.e. as droplet removal efficiency increases, allowable gas throughput decreases.
Table 2-3 provides a summary of performance parameters.
where
Vt = Velocity, ft/s.
ρl = Density of liquid droplet, lb/ft3
ρg = Density of gas, lb/ft3
Qg = actual gas flow rate, ft3/sec
A = Filter area, ft2

The K capacity factor for mesh pads is given in Table 2-3 and correction in table 2-4.

Table 2-3. Mesh Pad Separation Performance.

Table 2-4. Adjustment of K Factor for Pressure.

Mesh pads normally operate efficiently over a range of 30–110% of the design gas rate.
The gas capacity of a wire-mesh pad is defined in terms of a K “constant” as given in Table 2-3.

Figure 2-39. Dimensions for the placement of a wire-mesh mist extractor. [H represents minimum height, and Hm must be at least 1 foot].

2.12.7.5: Micro-Fiber
Micro-fiber mist extractors use very small diameter fibers, usually less than 0.02 mm, to capture very small droplets. Gas and liquid flow is horizontal and co-current. Because the micro-fiber unit is manufactured from densely packed fiber, drainage by gravity inside the unit is limited.
Much of the liquid is eventually pushed through the micro-fiber and drains on the downstream face. The surface area of a micro-fiber mist extractor can be 3 to 150 times that of a wire-mesh unit of equal volume. (Refer to figure 2-40).

Fig. 2-40. Micro fiber mist extractor.

Table 2-5 Major parameters in mist extractor selection.

2.12.7.6: Other Configurations
Some separators use centrifugal mist extractors, discussed earlier in this chapter, that cause liquid droplets to be separated by centrifugal force (refer to Figure 2-41). These units can be more efficient than either wire-mesh or vanes and are the least susceptible to plugging.
However, they are not in common use in production operations because their removal efficiencies are sensitive to small changes in flow. In addition, they require relatively large pressure drops to create the centrifugal force. To a lesser extent, random packing is sometimes used for mist extraction, as shown in Figure 2-42. The packing acts as a coalescer.

2.12.7.7: Final Selection
The selection of a type of mist extractor involves a typical cost-benefit analysis. Wire-mesh pads are the cheapest, but mesh pads are the most susceptible to plugging with paraffins, gas hydrates, etc. With age, mesh pads also tend to deteriorate and release wires and/or chunks of the pad into the gas stream. This can be extremely damaging to downstream equipment, such as compressors. Vane units, on the other hand, are more expensive. Typically, vane units are less susceptible to plugging and deterioration than mesh pads. Micro-fiber units are the most expensive and are capable of capturing very small droplets but, like wire mesh pads, are susceptible to plugging. The selection of a type of mist extractor is affected by the fluid characteristics, the system requirements, and the cost.

Fig. 2-41. Vertical separator equipped with centrifugal mist extractor.

Fig. 2-42. A coalescing pack mist extractor.

2.13: Control Components of Gas–Oil Separators
Gas–oil separators are generally equipped with the following control devices and internal components.
Liquid Level Controller
The liquid level controller (LLC) is used to maintain the liquid level inside the separator at a fixed height. In simple terms, it consists of a float that exists at the liquid–gas interface and sends a signal to an automatic valve on the oil outlet. The signal causes the valve to open or close, thus allowing more or less liquid out of the separator to maintain its level inside the separator.

Pressure Control Valve
The pressure control valve (PCV) is an automatic backpressure valve that exists on the gas stream outlet. The valve is set at a prescribed pressure.
It will automatically open or close, allowing more or less gas to flow out of the separator to maintain a fixed pressure inside the separator.

Pressure Relief Valve
The pressure relief valve (PRV) is a safety device that will automatically open to vent the separator if the pressure inside the separator exceeded the design safe limit.

Shut down valves
Shut down valves are usually installed at the inlet of separator to protect the vessel by preventing the incoming flow in case of vessel high pressure or high liquid level. Also it is usually installed at the outlet lines to prevent the flow out in case of very low liquid level or very low pressure.

yasserkassem · Post by **yasserkassem** » Tue Mar 17, 2020 9:36 pm

Two-phase Oil and Gas Separation -Chapter 2 - part 2

2.14: Operating Problems
2.14.1: Foamy Crude
The major cause of foam is the presence of impurities other than water in crude. One impurity that always causes foam is CO2. Workover fluids sometimes may be incompatible with the wellbore fluids, and will cause foam. Foam presents no problem within a separator if the internal design assures adequate time or sufficient coalescing surface for the foam to “break.”
Foaming in a separating vessel is a problem due to:
1. Foam will occupy a large space in the separator that otherwise would be available for the separation process; therefore, the separator efficiency will be reduced.
2. The foam will disrupt the operation of the level controller, since it has a density between that of the liquid and gas.
3. In case of existence of a foam bank, it will be possible for some of the foam to escape with gas outlet or with liquid outlet. Causing a problem in both cases.

The foaming tendencies of any oil can be determined with laboratory tests. One of the tests is ASTM D 892, which involves bubbling air through the oil. Alternatively, the oil may be saturated with its associated gas and then expanded in a gas container. This alternative test more closely models the actual separation process. Both of these tests are qualitative. There is no standard method of measuring the amount of foam produced or the difficulty in breaking the foam. Foaming is not possible to predict ahead of time without laboratory tests. However, foaming can be expected where CO2 is present in small quantities (1–2%). It should be noted that the amount of foam is dependent on the pressure drop to which the inlet liquid is subjected, as well as the characteristics of the liquid at separator conditions.
Comparison of foaming tendencies of a known oil to a new one, about which no operational information is known, provides an understanding of the relative foam problem that may be expected with the new oil as weighed against the known oil.
The effects of temperature on a foamy oil are interesting. Changing the temperature at which a foamy oil is separated has two effects on the foam. The first effect is to change the oil viscosity. That is, an increase in temperature will decrease the oil viscosity, making it easier for the gas to escape from the oil. The second effect is to change the gas-oil equilibrium. A temperature increase will increase the amount of gas, which evolves from the oil.
It’s very difficult to predict the effects of temperature on the foaming tendencies of an oil. However, some general observations have been made. For low API gravity crude (heavy oils) with low GORs, increasing the operating temperature decreases the oils’ foaming tendencies. Similarly, for high API crude (light oils) with high GORs, increasing the operating temperature decreases the oils’ foaming tendencies. However, increasing the operating temperature for a high API gravity crude (light oil) with low GORs may increase the foaming tendencies. Oils in the last category are typically rich in intermediates, which have a tendency to evolve to the gas phase as the temperature increases. Accordingly, increasing the operating temperature significantly increases gas evolution, which in turn increases the foaming tendencies.
Foam depressant chemicals often will do a good job in increasing the capacity of a given separator. However, in sizing a separator to handle a specific crude, the use of an effective depressant should not be assumed because characteristics of the crude and of the foam may change during the life of the field. Also, the cost of foam depressants for high-rate production may be prohibitive. Sufficient capacity should be provided in the separator to handle the anticipated production without use of a foam depressant or inhibitor.

2.14.2: Paraffin
Separator operation can be adversely affected by an accumulation of paraffin. Coalescing plates in the liquid section and mesh pad mist extractors in the gas section are particularly prone to plugging by accumulations of paraffin. Where it is determined that paraffin is an actual or potential problem, the use of plate-type or centrifugal mist extractors should be considered. Manways, handholes, and nozzles should be provided to allow steam, solvent, or other types of cleaning of the separator internals.
The bulk temperature of the liquid should always be kept above the cloud point of the crude oil.

2.14.3: Sand
Accumulation of san in the bottom of separators is serious operation problem, causing separator size reduction, cutout of valve trim, and plugging of separator internals. Accumulations of sand can be removed by periodically injecting water or steam in the bottom of the vessel so as to suspend the sand during draining. Figure 2-26 is a cutaway of a sand wash and drain system fitted into a horizontal separator fitted with sand jets and an inverted trough.
Sometimes a vertical separator is fitted with a cone bottom. This design would be used if sand production was anticipated to be a major problem.
The cone is normally at an angle of between 450 and 600 to the horizontal.
If a cone is installed, it could be part of the pressure-containing walls of the vessel (refer to Figure 2-43), or for structural reasons, it could be installed internal to the vessel cylinder (refer to Figure 2-43). In such a case, a gas equalizing line must be installed to assure that the vapor behind the cone is always in pressure equilibrium with the vapor space.
Plugging of the separator internals is a problem that must be considered in the design of the separator. A design that will promote good separation and have a minimum of traps for sand accumulation may be difficult to attain, since the design that provides the best mechanism for separating the gas, oil, and water phases probably will also provide areas for sand accumulation. A practical balance for these factors is the best solution.

2.14.4: Gas Blowby
Gas blowby occurs when free gas escapes with the liquid phase and can be an indication of low liquid level, vortexing, or level control failure. This could lead to a very dangerous situation. If there is a level control failure and the liquid dump valve is open, the gas entering the vessel will exit the liquid outlet line and would have to be handled by the next downstream vessel in the process. Unless the downstream vessel is designed for the gas blowby condition, it can be over-pressured. Gas blowby can usually be prevented by installing a level safety low sensor (LSL) that shuts in the inflow and/or outflow to the vessel when the liquid level drops to 10–15% below the lowest operating level. In addition, downstream process components should be equipped with a pressure safety high (PSH) sensor and a pressure safety valve (PSV) sized for gas blowby.

2.14.5: Liquid Carryover
Liquid carryover occurs when free liquid escapes with the gas phase and can indicate high liquid level, damage to vessel internals, foam, improper design, plugged liquid outlets, or a flow rate that exceeds the vessel’s design rate. Liquid carryover can usually be prevented by installing a level safety high (LSH) sensor that shuts in the inlet flow to the separator when the liquid level exceeds the normal maximum liquid level by some percentage, usually 10–15%.

Fig. 2- 43. Vertical separator with a pressure containing cone bottom, and vertical separator fitted with an internal cone bottom and an equalizing line.

2.14.6: Liquid Slugs
Two-phase flow lines and pipelines tend to accumulate liquids in low spots in the lines. When the level of liquid in these low spots rises high enough to block the gas flow, then the gas will push the liquid along the line as a slug. Depending on the flow rates, flow properties, length and diameter of the flow line, and the elevation change involved, these liquid slugs may contain large liquid volumes.
Situations in which liquid slugs may occur should be identified prior to the design of a separator. The normal operating level and the high-level shutdown on the vessel must be spaced far enough apart to accommodate the anticipated slug volume. If sufficient vessel volume is not provided, then the liquid slugs will trip the high-level shutdown.
When liquid slugs are anticipated, slug volume for design purposes must be established. Then the separator may be sized for liquid flow-rate capacity using the normal operating level. The location of the high-level set point may be established to provide the slug volume between the normal level and the high level. The separator size must then be checked to ensure that sufficient gas capacity is provided even when the liquid is at the high-level set point. This check of gas capacity is particularly important for horizontal separators because, as the liquid level rises, the gas capacity is decreased. For vertical separators, sizing is easier as sufficient height for the slug volume may be added to the vessel’s seam-to-seam length.
Often the potential size of the slug is so great that it is beneficial to install a large pipe volume upstream of the separator. The geometry of these pipes is such that they operate normally empty of liquid, but fill with liquid when the slug enters the system. This is the most common type of “slug catcher” used when two-phase pipelines are routinely pigged.
Figure 2-14 is a schematic of a liquid finger slug catcher.
2.15: Stage Separation
2.15.1: Initial Separation Pressure
Because of the multicomponent nature of the produced fluid, the higher the pressure at which the initial separation occurs, the more liquid will be obtained in the separator. This liquid contains some light components that vaporize in the stock tank downstream of the separator. If the pressure for initial separation is too high, too many light components will stay in the liquid phase at the separator and be lost to the gas phase at the tank. If the pressure is too low, not as many of these light components will be stabilized into the liquid at the separator and they will be lost to the gas phase.
This phenomenon, which can be calculated using flash equilibrium techniques discussed in previous chapter, is shown in Figures 2-44 and 2-45.

Fig. 2-44. Single stage separation.

It is important to understand this phenomenon qualitatively. The tendency of any one component in the process stream to flash to the vapor phase depends on its partial pressure. The partial pressure of a component in a vessel is defined as the number of molecules of that component in the vapor space divided by the total number of molecules of all components in the vapor space times the pressure in the vessel [refer to Eq. (2-3)]:
PPN =P × MolesN / ∑ MolesN Eq. 2-3
where
PPN = partial pressure of component “N,”
MolesN = number of moles of component “N,”
Ʃ MolesN = total number of moles of all components,
P = pressure in the vessel, psia.
Thus, if the pressure in the vessel is high, the partial pressure for the component will be relatively high and the molecules of that component will tend toward the liquid phase. This is seen by the top line in Figure 2-45.
As the separator pressure is increased, the liquid flow rate out of the separator increases.
The problem with this is that many of these molecules are the lighter hydrocarbons (methane, ethane, and propane), which have a strong tendency to flash to the gas state at stock-tank conditions (atmospheric pressure). In the stock tank, the presence of these large numbers of molecules creates a low partial pressure for the intermediate-range hydrocarbons (butanes, pentane, and heptane) whose flashing tendency at stock tank conditions is very susceptible to small changes in partial pressure. Thus, by keeping the lighter molecules in the feed to the stock tank, we manage to capture a small amount of them as liquids, but we lose to the gas phase many more of the intermediate-range molecules. That is why beyond some optimum point there is actually a decrease in stock-tank liquids by increasing the separator operating pressure.

2.15.2: Stage Separation
Figure 2-44 deals with a simple single-stage process. That is, the fluids are flashed in an initial separator and then the liquids from that separator are flashed again at the stock tank. Traditionally, the stock tank is not normally considered a separate stage of separation, though it most assuredly is.
Figure 2-46 shows a three-stage separation process. The liquid is first flashed at an initial pressure and then flashed at successively lower pressures two times before entering the stock tank.
Because of the multicomponent nature of the produced fluid, it can be shown by flash calculations that the more stages of separation after the initial separation, the more light components will be stabilized into the liquid phase. This can be understood qualitatively by realizing that in a stage separation process the light hydrocarbon molecules that flash are removed at relatively high pressure, keeping the partial pressure of the intermediate hydrocarbons lower at each stage. As the number of stages approaches infinity, the lighter molecules are removed as soon as they are formed and the partial pressure of the intermediate components is maximized at each stage. The compressor horsepower required is also reduced by stage separation as some of the gas is captured at a higher pressure than would otherwise have occurred. This is demonstrated by the example in Table 2-6.

Table. 2-6. Effect of separation pressure for a rich condensate stream.

Fig. 2-45. Effect of separator pressure on liquid recovery.

Fig. 2-46. Stage separation
2.15.3: Selection of Stages
As shown in Figure 2-47, as more stages are added to the process there is less and less incremental liquid recovery. The diminishing income for adding a stage must more than offset the cost of the additional separator, piping, controls, space, and compressor complexities. It is clear that for each facility there is an optimum number of stages. In most cases, the optimum number of stages is very difficult to determine as it may be different from well to well and it may change as the well’s flowing pressure declines with time. Table 2-7 is an approximate guide to the number of stages in separation, excluding the stock tank, which field experience indicates is somewhat near optimum. Table 2-7 is meant as a guide and should not replace flash calculations, engineering studies, and engineering judgment.

Fig.2-47. Incremental liquid recovery versus number of separator stages.

Table. 2-7. Stage separation guidelines.

2.15.4: Fields with Different Flowing Tubing Pressures
The discussion to this point has focused on a situation where all the wells in a field produce at roughly the same flowing tubing pressure, and stage separation is used to maximize liquid production and minimize compressor horsepower. Often, stage separation is used because different wells producing to the facility have different flowing tubing pressures. This could be because they are completed in different reservoirs, or are located in the same reservoir but have different water production rates. By using a manifold arrangement and different primary separator operating pressures, there is not only the benefit of stage separation of high-pressure liquids, but also conservation of reservoir energy. High-pressure wells can continue to flow at sales pressure requiring no compression, while those with lower tubing pressures can flow into whichever system minimizes compression.

2.15.5: Determining Separator Operating Pressures
The choice of separator operating pressures in a multistage system is large. For large facilities many options should be investigated before a final choice is made. For facilities handling less than 50,000 bpd, there are practical constraints that help limit the options.
A minimum pressure for the lowest-pressure stage would be in the 25- to 50-psig range. This pressure will probably be needed to allow the oil to be dumped to a treater or tank and the water to be dumped to the water treating system. The higher the operating pressure, the smaller the compressor needed to compress the flash gas to sales. Compressor horsepower requirements are a function of the absolute discharge pressure divided by the absolute suction pressure.
Increasing the low-pressure separator pressure from 50 psig to 200 psig may decrease the compression horsepower required by 33%. However, it may also add backpressure to wells, restricting their flow, and allow more gas to be vented to atmosphere at the tank. Usually, an operating pressure of between 50 and 100 psig is optimum.

As stated before, the operating pressure of the highest-pressure separator will be no higher than the sales gas pressure. A possible exception to this could occur where the gas lift pressure is higher than the sales gas pressure. In choosing the operating pressures of the intermediate stages, it is useful to remember that the gas from these stages must be compressed.
Normally, this will be done in a multistage compressor. For practical reasons, the choice of separator operating pressures should match closely and be slightly greater than the compressor inter-stage pressures.

Fig. 2-48. Compressor stages and inlet points of separated gas from multistage separation.

The most efficient compressor sizing will be with a constant compressor ratio per stage. Therefore, an approximation of the intermediate separator operating pressures can be derived from

R = (Pd/Ps)1/n Eq. 2-4

where
R = Compression ratio per stage,
Pd = discharge pressure, psia,
Ps = suction pressure, psia,
n = number of stages.
In order to minimize inter-stage temperatures, the maximum ratio per stage will normally be in the range of 3.6 to 4.0. That means that most production facilities will have either two- or three-stage compressors. A two-stage compressor only allows for one possible intermediate separator operating pressure. A three-stage allows for either one operating at second- or third-stage suction pressure or two intermediate separators each operating at one of the two compressor intermediate suction pressures.( fig. 2-48).

2.15.6: Two-Phase vs. Three-Phase Separators
In our example process the high- and intermediate-stage separators are two-phase, while the low-pressure separator is three-phase. This is called a “free-water knockout” (FWKO) because it is designed to separate the free water from the oil and emulsion, as well as separate gas from liquid.
The choice depends on the expected flowing characteristics of the wells.
If large amounts of water are expected with the high-pressure wells, it is possible that the size of the other separators could be reduced if the high-pressure separator was three-phase.
2.16: Separator calculation basics.
2.16.1: Liquid Handling and Liquid Retention Time
To assure that the liquid and gas reach equilibrium at separator pressure, a certain liquid storage is required. This is defined as “retention time” or “residence time”, or the average time a molecule of liquid is retained in the vessel, assuming plug flow. The retention time is thus the volume of the liquid storage in the vessel divided by the liquid flow rate.
The design criterion for two phase separator liquid handling capacity is typically based on:
• Liquid degassing requirements.
• Process control/stability requirements.
Generally, one or the other of these factors will dictate. Liquid capacity is typically specified in terms of residence time, which must be translated into vessel layout requirements for dimensioning purposes. Residence time establishes the separator volume required for the liquid as shown in Eq 2-5 :

V = (W (t))/1440 Eq. 2-5
Where
W = Liquid handling capacity, bbl/day.
V = Liquid settling volume, bbl ( bbl = 5.615 ft3 )
t = Retention time, minutes

For most applications retention times between 30 s and 3 min have been found to be sufficient. Where foaming crude is present, retention times up to four times this amount may be needed. In the absence of liquid or laboratory data, the guidelines presented in Table 2-8 can be used.

Table 2-8, API 12J Liquid retention time for gas oil separators.

Values represented in table 2-8, are based on liquid degassing requirements. In practice, process control stability and operability requirements will often override the degassing requirements. The retention time requirements given in table 2-8, is not specific to vessel orientation. However, the liquid degassing process actually involves the separation of gas bubbles from the liquid phase, which under ideal conditions can be described by the gravity settling equation.
A gas bubble size of 150–200 microns has been suggested by several sources for calculating vessel liquid handling requirements for a degassing constraint according to gravity settling theory.
Example 2-1
Calculate the retention time for horizontal separator 2 ft. diameter and 5 ft. length.
The liquid flow rate is 2000 bbl/day, and is operating half full of liquid.
Solution:
W = 2000 bbl/day.
Liquid settling volume V (50% of vessel) =(π D2 L )/(4 x 2)
V = 3.14 x 4 x5 /8 = 7.85 ft3. = 7.85 x 0.178 = 1.4 bbl.
From eq. 2-6
1.4 = (2000 (t))/1440
t = 1 minute.

Example 2-2
Calculate the retention time for vertical separator 2 ft. diameter and 6 ft. height.
The liquid flow rate is 1000 bbl/day, and the liquid level is 30% height.
Solution:
W = 1000 bbl/day.
Liquid settling volume V (30% of vessel) =(π D2 h ) x 30/(4 x 100)
V = 3.14 x 4 x6 x 30 /(4 x 100) = 5.65 ft3. = 5.65 x 0.178 = 1.0 bbl.
From eq. 2-6
1.0 = (1000 (t))/1440
t = 1.4 minutes.

2.16.2: Gas retention time
The separator must have a sufficient area for gas flow in which the gas will travel in a certain time more than the time required for the liquid droplets settle to the liquid accumulation section.
The gas retention time is calculated from equation 2-6.
Retention time Seconds= (V/Q) Eq. 2-6

where
V is the volume of vessel for gas flow ft3
Q is the gas flow rate at operation conditions ft3/s.

2.16.3: Gas velocity
The gas velocity is determined by the flow rate of gas (Q) ft3/s at operating conditions divided by the cross sectional area where the gas is flowing.
Vg = (Q) ft3/s /(A) ft2. = ft/s Eq. 2-7
Example 2-3
For horizontal separator 2 ft. diameter and 5 ft. length, calculate the gas retention time and gas velocity for gas flow rate 8 MMscfd, operating pressure 300 psia, temperature 80 0F. Compressibility factor 0.95.
Liquid level 50%.
Solution:
Volume of vessel for gas flow is 50% of vessel (As example 2-1) = 7.85 ft3
Area of gas flow = =(π D2 ) * 50 /(4*100 ) = 1.57 ft2 ( In case of horizontal vessel 50%of area will be for gas flow, and 50% for liquid (assuming the vessel is operated 50% liquid capacity, while area for gas flow in vertical separator is 100% of the top part of the vessel “ over liquid level”.)
Temperature, 0R = 80+460 = 540
Pressure, psia = 300
Compressibility factor Z = 0.95 (given).
Flow rate = 8 MMscfd
Remember ( R =10.73. ) and (n = flowing gas at standard conditions,( ft3 /379.5 ))
From eq. 1-10, PV= nzRT
Rate of flowing gas at operating condition =
Q = (8 x 106 x 0.95 x 10.73 x 540)/ (300 x 379) ft3/day
Q = 387299 ft3/day
Q = 4.48 ft3/s
Gas retention time t = Volume of gas room / flow rate = 7.85/4.48
Gas retention time = 1.74 seconds.
Gas velocity = 4.48/1.57 = 2.9 ft/s

Note that in the previous example we assumed the effective length for the separator is the same as the seam to seam length given, while sometimes the gas outlet is not in the end of the vessel, so the effective length for gas flow is less than the seam to seam length.

Example 2-4
For vertical separator 2 ft. diameter and 6 ft. height, the liquid level is 30% height.
Calculate the gas retention time and gas velocity for gas flow rate 8 MMscfd, operating pressure 300 psia, temperature 80 0F. Compressibility factor 0.95.
Solution:
Volume of vessel for gas flow is 70% of vessel = (π D2 h) x 70 / (4 x 100)
V = 3.14 x 4 x6 x 70 / (4 x 100) = 13.2 ft3
Area of gas flow = (π D2 ) /4 = 3.14 ft2
(Flow of gas is through full diameter in case of vertical separator.)
Rate of flowing gas at operating condition =
Q = (8 x 106 x 0.95 x 10.73 x 540)/ (300 x 379) ft3/day
Q = 387299 ft3/day
Q = 4.48 ft3/s
Gas retention time t = Volume of flow area/flow rate = 13.2/4.48
Gas retention time = 2.95 seconds.
Gas velocity = 4.48/3.14 = 1.42 ft/s

From previous examples, it’s clear that vertical separators have low liquid capacity than the same diameter horizontal separator, while it can handle more gas than horizontal vessel at same gas velocity.

2.16.4: Liquid Re-entrainment
Liquid re-entrainment is a phenomenon caused by high gas velocity at the gas-liquid interface of a separator. Momentum transfer from the gas to the liquid causes waves and ripples in the liquid, and then droplets are broken away from the liquid phase.
The general rule of thumb that calls for limiting the slenderness ratio to a maximum of 4 or 5 is applicable for half-full horizontal separators. “The flow entry in vertical separators is high enough from liquid level, and there is no gas flow near the gas liquid interface as in horizontal separators”.
Liquid re-entrainment should be particularly considered for high-pressure separators sized on gas-capacity constraints. It is more likely at higher operating pressures >1,000 psig and higher oil viscosities (<300API).

2.16.5: Droplet Size (Liquid in gas phase)
The purpose of the gravity settling section of the vessel is to condition the gas for final polishing by the mist extractor. To apply the settling equations to separator sizing, a liquid droplet size to be removed must be selected. From field experience, it appears that if 140-micron droplets are removed in this section, the mist extractor will not become flooded and will be able to perform its job of removing those droplets between 10- and 140-micron diameters. The gas capacity design equations in this section are all based on 140-micron removal. In some cases, this will give an overly conservative solution. The techniques used here can be easily modified for any droplet size.
In this section we are addressing separators used in oil field facilities. These vessels usually require a gravity settling section design for removal of droplet 140-micron in size. There are special cases where the separator is designed to remove only very small quantities of liquid that could condense due to temperature or pressure changes in a stream of gas that has already passed through a separator and a mist extractor. These separators, commonly called “gas scrubbers,” could be designed for removal of droplets on the order of 500 microns without fear of flooding their mist extractors. Fuel gas scrubbers, compressor suction scrubbers, and contact tower inlet scrubbers are examples of vessels to which this might apply.
Flare or vent scrubbers are designed to keep large slugs of liquid from entering the atmosphere through the vent or relief systems. In vent systems the gas is discharged directly to the atmosphere, and it is common to design the scrubbers for removal of 300- to 500-micron droplets in the gravity settling section. A mist extractor is not included because of the possibility that it might get plugged, thus creating a safety hazard.
In flare systems, where the gas is discharged through a flame, there is the possibility that burning liquid droplets could fall to the ground before being consumed. It is still common to size the gravity settling section for 300- to 500-micron removal, which the API guideline for refinery flares indicates is adequate to ensure against a falling flame. In critical locations, such as offshore platforms, many operators include a mist extractor as an extra precaution against a falling flame. If a mist extractor is used, it is necessary to provide safety relief protection around the mist extractor in the event that it becomes plugged.

2.17: Design Principles and sizing of Oil-gas Separator
Since the drag force is one of the main factors affecting liquid droplet in gas phase settling velocity and consequently vessel design parameters, we will present two methods of calculation the drag coefficient and calculation of the settling velocity.
The first method assumes a drag coefficient value and proceed calculations to check the accuracy of the assumed value, and then repeat if the value is not accepted, and repeat until reaching satisfied value.
The second method uses the drag coefficient value extracted from chart, or calculate it with equations.
Each method will be followed by examples.

2.17.1: First method Design Theory
2.17.1.1: Settling
In the gravity settling section of a separator, liquid droplets are removed using the force of gravity. Liquid droplets, contained in the gas, settle at a terminal or “settling” velocity.
If the flow around the droplet were laminar (Re < 1), then Stokes’ law would govern and The drag coefficient CD will equal

CD = 24 / Re Eq.2-8

Where
CD =drag coefficient,
Re = Reynolds number, which is dimensionless.
It can be shown that in such a gas the droplet settling velocity would be given by:

Vt = 1.78 x 10-6 (ΔSG) d2m / µ Eq. 2-9

ΔSG = specific gravity difference (Kg/l).
Vt =terminal (settling velocity) of the droplet, ft/s,
dm =droplet diameter, microns,
µ =viscosity of the gas, cp.

Unfortunately, for production facility designs it can be shown that Stokes’ law does not govern, and the following more complete formula for drag coefficient must be used (refer to Figure 2-49)

Fig. 2-49. Reynolds number and drag coefficient.

The drag coefficient will equal

CD = (24 /Re) + (3 / Re0.5) + 0.34 Eq. 2-10

Equating drag and buoyant forces, the terminal settling velocity is given by

Vt = 0.0119 [(ρd – ρc ) dm / CD ρc ]0.5 Eq. 2-11
where
ρd =density of the liquid droplet, lb/ft3,
ρc = density of the continuous phase (medium) where droplet will travel. “gas in case of liquid droplet settle from gas phase”, lb/ft3,
Vt =terminal (settling velocity) of the droplet, ft/s,
dm =droplet diameter, microns,

Equations 2-10, and 2-11, can be solved by an iterative process. Start by assuming a value of CD, such as 0.34, and solve Eq. (2-11) for Vt .Then, using Vt , solve for Re. using equation 2-12, or 2-13.
Re = 0.0049 dmVt ρc / µ Eq. 2-12
Or
Re = 1488 DmVt ρc / µ Eq. 2-13
where
ρc = density of the continuous phase (medium) where droplet will travel “gas in case of liquid droplet settle from gas phase”, lb/ft3,
Vt =terminal (settling velocity) of the droplet, ft/s,
Dm = droplet diameter, ft.
dm =droplet diameter, micron
µ =viscosity of the gas, cp.

Then, Eq. (2-10) may be solved for CD. If the calculated value of CD equals the assumed value, the solution has been reached. If not, then the procedure should be repeated using the calculated CD as a new assumption. The original assumption of 0.34 for CD was used because this is the limiting value for large Reynolds numbers.
The iterative steps are shown below:
Start with
Vt = 0.0119 [(ρd – ρc ) dm /0.34 x ρc ]0.5
Calculate
Re = 0.0049 dmVt ρc / µ (dm) in micron (Or)
Re = 1488 DmVt ρc / µ (Dm) in ft
From Re calculate CD as follows:
CD = (24/Re) +[ 3/(Re)0.5 ] + 0.34
Recalculate Vt using
Vt = 0.0119 [(ρd – ρc ) dm / CD ρc ]0.5
Go to step 2 and iterate

2.17.1.2: Separator Design (Horizontal Separators Sizing)
The guidelines presented in this section can be used for the initial sizing of a horizontal separator 50% full of liquid.
When sizing a horizontal separator, it is necessary to choose a seam-to seam vessel length and a diameter. This choice must satisfy:
1- The conditions for gas capacity that allow the liquid droplets to fall from the gas to the liquid volume as the gas traverses the effective length of the vessel.
2- Provide sufficient retention time to allow the liquid to reach equilibrium.
Figure 2-50 shows a vessel 50% full of liquid, which is the model used to develop sizing equations for a horizontal separator.

Fig. 2-50. Model of a horizontal separator.

2.17.1.3: Gas Capacity Constraint
The principles of liquid droplets settling through a gas can be used to develop an equation to size a separator for a gas flow rate. The gas capacity constraint equations are based on setting the gas retention time equal to the time required for a droplet to settle to the liquid interface. For a vessel 50% full of liquid, and separation of liquid droplets from the gas, the following equation may be derived:

dLeff = 420 (TZQ/P) [CD ρc/(ρd- ρc) dm ]0.5 Eq. 2-14
where
d =vessel internal diameter, in,
Leff =effective length of the vessel where separation occurs, ft,
T =operating temperature, 0R,
Qg =gas flow rate, MMscfd,
P =operating pressure, psia,
Z =gas compressibility,
CD =drag coefficient,
dm =liquid droplet to be separated, micron,
ρc=density of continuous phase “medium” (gas in this case), lb/ft3,
ρd=density of droplet, lb/ft3.

2.17.1.4: Liquid Capacity Constraint
Two-phase separators must be sized to provide some liquid retention time so the liquid can reach phase equilibrium with the gas (evolving gas bubbles). For a vessel 50% full of liquid, with a specified liquid flow rate and retention time, the following may be used to determine vessel size.

d2 Leff = tr Ql / 0.7 Eq. 2-15

where
tr = desired retention time for the liquid, min,
Ql = liquid flow rate, bpd.

2.17.1.5: Seam-to-Seam Length
The effective length may be calculated from previous equations Eq. (2-14, and 2-15). From this, a vessel seam-to-seam length may be determined. The actual required seam-to-seam length is dependent on the physical design of the internals of the vessel.

Figure 2-51. Approximate seam-to-seam length of a horizontal separator one-half full.

As shown in Figure 2-51, for vessels sized on a gas capacity basis, some portion of the vessel length is required to distribute the flow evenly near the inlet diverter. Another portion of the vessel length is required for the mist extractor. The length of the vessel between the inlet diverter and the mist extractor with evenly distributed flow is the Leff calculated from Eq. (2-14, and 2-15).
Based on these concepts coupled with field experience, the seam-to-seam length of a vessel may be estimated as the larger of the following.

Lss = Leff + d/12 (For gas capacity) Eq. 2-16

For vessels sized on a liquid capacity basis, some portion of the vessel length is required for inlet diverter flow distribution and liquid outlet.
The seam-to-seam length is estimated as follow:

Lss = (4/3) Leff (For liquid capacity) Eq. 2-17a

For diameters => 36 in. Lss = ( Leff + 2.5 ) Eq. 2-17b

2.17.1.6: Slenderness Ratio
Equations (2-14) and (2-15) allow for various choices of diameter and length. For each vessel design, a combination of Leff and d exists that will minimize the cost of the vessel. It can be shown that the smaller the diameter, the less the vessel will weigh and thus the lower its cost. There is a point, however, where decreasing the diameter increases the possibility that high velocity in the gas flow will create waves and re-entrain liquids at the gas-liquid interface. Experience has shown that if the gas capacity governs and the length divided by the diameter, referred to as the “slenderness ratio,” is greater than 4 or 5, re-entrainment could become a problem. Most two-phase separators are designed for slenderness ratios between 3 and 5.
Alternatively, instead of calculating the slenderness ratio, and check if it is accepted or not, the user may select a combination of diameter and length from standard separator sizes.
The following two tables are included in API 12J.

Table 2-9, API 12J Standard horizontal separator sizes.

2.17.1.7: Procedure for Sizing Horizontal Separators—Half Full
1. The first step in sizing a horizontal separator is to establish the design basis. This includes specifying the maximum and minimum flow rates, operating pressure and temperature, droplet size to be removed, etc.
2. Prepare a table with calculated values of Leff for selected values of d that satisfy Eq. (2-14) ”the gas capacity constraint”.
Calculate Lss using Eq. (2-16).

dLeff = 420 (TZQg/P) [CD ρc/(ρd- ρc) dm ]0.5 Eq. 2-14

Lss = Leff + d/12 for gas capacity Eq. 2-16

3. For the same values of d, calculate values of Leff using Eq. (2-15) for liquid capacity and list these values in the same table.
Calculate Lss using Eq. (2-17a or b).

d2 Leff = tr Ql / 0.7 Eq. 2-15

Lss = (4/3) Leff or Lss = ( Leff + 2.5 ) Eq. 2-17a & 2-17b

4. For each d, the larger Leff should be used.(The larger number will cover both constraints).
5. Calculate the slenderness ratio, 12Lss/do, and list for each d. Select a combination of d and Lss that has a slenderness ratio between 3 and 5. Lower ratios can be chosen if dictated by available space, but they will probably be more expensive. (or select a suitable standard separator size using table 2-9).
6. When making a final selection, it is always more economical to select a standard vessel size. Vessels with outside diameters up through 24 inches have nominal pipe dimensions. Vessels with outside diameters larger than 24 inches are typically rolled from plate with diameter increments of 6 inches. The shell seam-to-seam length is expanded in 2.5-ft segments and is usually from 5 ft to 10 ft. Standard separator vessel sizes from API 12J are listed in table 2-9.

2.17.1.8: Procedure Vertical Separators’ Sizing
The guidelines presented in this section can be used for initial sizing of a vertical two-phase separator.
In vertical separators, a minimum diameter must be maintained to allow liquid droplets to separate from the vertically moving gas. The liquid retention time requirement specifies a combination of diameter and liquid volume height. Any diameter greater than the minimum required for gas capacity can be chosen. Figure 2-52 shows the model used for a vertical separator.

2.17.1.9: Gas Capacity Constraint
The principles of liquid droplets settling through a gas can be used to develop an equation to size a separator for a gas flow rate. By setting the gas retention time equal to the time required for a droplet to settle to the liquid interface, the following equation may be derived.

d2 = 5040 (TZQg/P) [CD ρc/(ρd- ρc) dm ]0.5 Eq. 2-18
where
d =vessel internal diameter, in.,
T =operating temperature, 0R,
Qg =gas flow rate, MMscfd,
P =operating pressure, psia,
Z =gas compressibility,
CD =drag coefficient,
dm =liquid droplet to be separated, micron,
ρc=density of continuous phase “ gas in this case”, lb/ft3,
ρl=density of droplet “liquid in this case”, lb/ft3.

1.17.1.10: Liquid Capacity Constraint
Two-phase separators must be sized to provide some liquid retention time so the liquid can reach phase equilibrium with the gas. For a specified liquid flow rate and retention time, the following may be used to determine a vessel size.
d2 h = tr Ql / 0.12 Eq. 2-19
where
h = height of the liquid volume, in.

Fig. 2-52. Model of a vertical separator.

2.17.1.11: Seam-to-Seam Length
As with horizontal separators, the specific design of the vessel internals will affect the seam-to-seam length. The seam-to-seam length of vertical vessels may be estimated based on the diameter and liquid height. As shown in Figure 2-53, allowance must be made for the gas separation section and mist extractor and for any space below the water outlet. The following equations may be used to estimate Lss.
Lss = (h+76)/12 for diameters <=36 in. Eq. 2-20

Lss = (h+d+40)/12 for diameters >36 in. Eq. 2-21

where
h = height of liquid level, in.,
d = vessel ID, in.
The larger of the Lss values from Eqs. (2-20 and 2-21) should be used.

2.17.1.12: Slenderness Ratio
As with horizontal separators, the larger the slenderness ratio, the less expensive the vessel will be. In vertical separators whose sizing is liquid dominated, it is common to choose slenderness ratios no greater than 4 to keep the height of the liquid collection section to a reasonable level.
Choices of between 3 and 4 are common, although height restrictions may force the choice of a lower slenderness ratio.
Alternatively, instead of calculating the slenderness ratio, and check if it is accepted or not, the user may select a combination of diameter and length from standard separator sizes.
The following two tables are included in API 12J.

Figure 2-53. Approximate seam-to-seam shell length for a vertical separator.

Table 2-10, API 12J Standard vertical separator sizes.

2-17-1-13: Procedure for Sizing Vertical Separators
1. The first step in sizing a vertical separator is to establish the design basis. This includes specifying the maximum and minimum flow rates, operating pressure and temperature, droplet size to be removed, etc.
2. Equation (2-18) may be used to determine the minimum required d. Any diameter larger than this value may be used.
d2 = 5040 (TZQg/P) [CD ρc/(ρd- ρc) dm ]0.5 Eq. 2-18
3. For a selected d, Eq. (2-19) may be used to determine h.
d2 h = tr Ql / 0.12 Eq. 2-19
4. From d and h, the seam-to-seam length may be estimated using Eq. (2-20) and (2-21). The larger value of Lss should be used.
Lss = (h+76)/12 for diameters <=36 in. Eq. 2-20

Lss = (h+d+40)/12 for diameters >36 in. Eq. 2-21

5. Check the slenderness ratio to determine if it is less than 4. (or select a suitable standard separator size using table 2-10).
6. When making a final selection, it is always more economical to select a standard vessel size. Vessels with outside diameters up through 24 inches have nominal pipe dimensions. Vessels with outside diameters larger than 24 inches are rolled from plate with diameter increments of 6 inches. The shell seam-to-seam length is expanded in 2.5-ft segments and is usually from 5 ft to 10 ft. Standard separator vessel sizes obtained from API 12J are listed in table 2-10.

Example 2-5: Sizing a Two Phase Vertical Separator
Given:
Gas flow rate: 10 MMscfd at 0.6 specific gravity
Oil flow rate: 2,000 BOPD at 40 0API
Operating pressure: 1,000 psia
Operating temperature: 600F
Droplet size removal: (dm) =140 microns
Retention time: 3 min
Solution:
Calculate CD.
ρd= 62.4 x [ 141.5/(131.5+40)] = 51.5 lb/ft3
From eq.1-16 ρg= 0.093 ((MW)P)/TZ lb/ft3
Z = 0.84 (from Chapter 1)
MW = 0.6 x 29 = 17.4
ρc (gas)= 0.093 x 17.4 x1,000/(520 x 0.84) = 3.7 lb/ft3
µ = 0.013 cp (from Chapter 1)
Assume CD = 0.34,
Vt = 0.0119 [(ρd – ρc ) dm / CD ρc ]0.5
Vt = 0.0119 [(51.5 – 3.7 ) x 140 / 0.34 x 3.7 ]0.5
Vt = 0.868 ft/s
Re = 0.0049 x 3.7 x 140 x 0.868 / 0.013 = 169.47
CD = (24/169.47) +[ 3/(169.47)0.5 ] + 0.34 = 0.712

Repeat using CD = 0.712
Vt = 0.599 ft/s
Re = 117
CD = 0.822

Repeat using CD = 0.822
Vt = 0.556
Re = 110
CD = 0.844

Repeat using CD = 0.844
Vt = 0.548
Re = 108
CD = 0.851

Repeat using CD = 0.851
Vt = 0.545
Re = 108
CD = 0.854 ok

Gas capacity constraint
d2 = 5040 (TZQg/P) [CD ρc/(ρd- ρc) dm ]0.5
d2 = 5040 (520 x 0.84 x 10/1000) x [0.851 x 3.71/(51.5 – 3.71) 140 ]0.5
d = 21.9 in.

Liquid capacity constraint

d2 h = tr Ql / 0.12
d2 h = 3 x 2000 / 0.12
d2 h = 50,000

Select different values of Diameter (d), larger than the minimum value of gas constraint, and check the height “h” in liquid constraint equation.

Assume D = 24 in.
h = 50,000 / (24)2
h = 86.8 in

Calculate Seam to Seam length

Lss = (h+76)/12 for diameters <=36 in.
Lss = (86.8+76)/12 in.
Lss = 13.6 ft

Calculate slenderness ratio =12Lss/d = 6.8
Slenderness ratio should be from 3 to 4, So select another larger diameter (30, 36, 42 in.) and repeat.

Assume d = 36 in.
h = 50,000 / (36)2
h = 38.6 in

Calculate Seam to Seam length
Lss = (h+76)/12 for diameters <=36 in.
Lss = (38.6+76)/12 in.
Lss = 9.55 ft

Calculate slenderness ratio =12Lss/d = 3.2
Slenderness ratio ok.

Selection will be 36 in. diameter, 10 ft. length.

Alternatively, after calculating the diameter 24”, and the seam to seam length, 13.6 ft., we can check in standard vertical separator table 2-10, we will find that the maximum seam to seam length for 24” diameter is 10ft. which will not match the required length (13.6). In this case a larger diameter is selected (36”) and proceed calculation to get the seam to seam length, where it found to be 9.55 ft. (i.e. 10 ft. from table 2-10).

Example 2-6: Sizing a Two Phase Horizontal Separator
Given:
Gas flow rate: 10 MMscfd at 0.6 specific gravity
Oil flow rate: 2,000 BOPD at 400API
Operating pressure: 1,000 psia
Operating temperature: 600F
Droplet size removal: 140 microns
Retention time: 3 minutes

Solution:
1. From example 2-5 we will use the calculated liquid and gas density, z factor, gas viscosity, and CD .
CD = 0.851

2. Gas capacity constraint
dLeff = 420 (TZQ/P) [CD ρc/(ρd- ρc) dm ]0.5
dLeff = 420 ((520 x 0.84 x 10)/1000) [0.851 x 3.71 /(51.5 – 3.71) 140 ]0.5
dLeff = 39.85

3. Liquid capacity constraint
d2 Leff = tr Ql / 0.7
d2 Leff = 3 x 2000 / 0.7
d2 Leff = 8571

4. Compute combinations of d and Lss satisfy the gas and liquid constraint.
5. Compute seam-to-seam length for various d (Table 2-11).
6. Compute slenderness ratios, 12Lss/d. Choices in the range of 3 to 4 are common.
Example, For d = 30 in. ( review solution steps in 2-17-1-7)
Gas Leff = 1.8 ft
Liquid Leff = 9.5 ft So liquid constraint, governs selection
Lss = Leff x (4/3)
Lss = 9.5 x (4/3) = 12.7
Slenderness ratios, 12Lss/d. = 12 x 12.7 / 30 = 5.1 (high, select a larger diameter).
For d = 36 in.
Gas Leff = 1.53 ft
Liquid Leff = 6.6 ft So liquid constraint, governs selection
Lss = (4/3) Leff For liquid capacity
Lss = 6.6 + 2.5 = 9.1 ( Eq. 2-17b)
Slenderness ratios, 12Lss/d. = 8.8 x 12 / 36 = 2.9 OK
Alternatively, after calculating the diameter 30”, and the seam to seam length, 12.7 ft, we can check in standard horizontal separator table 2-9. We will find that API accepted diameter 30” with 15 ft. length, In this case selection can be 30” – 15 ft. or 36” – 10 ft.( Slenderness ratio is higher than 5, but the liquid capacity governs the dimensions not the gas capacity, so there will not be possibility of re-entrainment. Review 2-17-1-6).

* Lss = Leff +2.5 governs
Table 2-11. Diameter and length combination for horizontal separator sizing.

2.17.2: Second method Design Theory
Liquid droplets will settle out of a gas phase if the gravitational force acting on the droplet is greater than the drag force of the gas flowing around the droplet. These forces can be described mathematically using the terminal or finite-settling velocity calculation, Eq 2-22.
Vt = [4gDm (ρd – ρc ) / (3 ρc CD)]0.5 Eq. 2-22
where
Vt = critical or terminal gas velocity necessary for particles of size Dm to drop or settle out of gas, ft/sec
g = acceleration due to gravity, 32.2 ft/sec2
Dm = droplet diameter, ft
Ρd = density of droplet or particle, lb/ft3
Ρc = density of continuous phase “gas in case of droplet of liquid settling from gas phase”, lb/ft3
CD = drag coefficient of particle, dimensionless

The Calculation may be proceeded by two methods, the first method is to extract the coefficient from graph, and the second method is to calculate the coefficient.

A - Getting drag coefficient from a graph.
The drag coefficient has been found to be a function of the shape of the particle and the Reynolds number of the flowing gas. For the purpose of this equation, particle shape is considered to be a solid, rigid sphere. The Reynolds number is defined as:
Re = 1488 DmVt ρc / µ Eq. 2-23
In this form, a trial and error solution is required (as proceeded in first method) since both particle size (Dm) and terminal velocity (Vt) are involved. To avoid trial and error, values of the drag coefficient are presented in fig. 2-54 as a function of the product of drag coefficient (CD) times the Reynolds number squared; this technique eliminates velocity from the expression. The abscissa of fig. 2-55 is given by:
CD (Re)2 = (0.95) (108) ρc (Dm)3 (ρd – ρc )/ µ2 Eq. 2-24

So [CD (Re)2] can be obtained from eq.2-24 and apply in fig. 2-54 to get the Drag coefficient (CD), And then apply the value is eq. 2-22

Figure 2-54. Drag coefficient.

B- Getting drag coefficient from different equations.
As an alternative to using eq. 2-24 and Fig. 2-54, the following approach is commonly used.
The curve shown in Fig. 2-49 can be simplified into three sections from which curve-fit approximations of the CD vs Re curve can be derived. When these expressions for CD vs Re are substituted into eq. 2-22, three settling laws are obtained as described below.

Stoke’s Law
At low Reynolds numbers (Fig. 2-49), a linear relationship exists between the drag coefficient and the Reynolds number (corresponding to laminar flow). Stoke’s Law applies in this case
Vt = 1488 g (Dm)2 (ρd – ρc ) / 18 µ Eq. 2-25

ρd =density of droplet fluid, lb/ft3, (= 62.4×SG)
ρc = density of the continuous phase, lb/ft3, (= 62.4×SG)
Dm =droplet diameter, ft.
µ =viscosity of the continuous phase, cp.
The droplet diameter corresponding to a Reynolds number of 2 was found to be less than 100 micron. For this reason, Stoke’s law is typically applicable for small droplet sizes and/or relatively high viscosity liquid phases.

Intermediate Law
For Reynold’s numbers between 2 and 500, the Intermediate Law applies, and the terminal settling law can be expressed as:
Vt = [3.49 g0.71 (Dm)1.14 (ρd – ρc)0.71 ] / [(ρc)0.29 µ0.43] Eq. 2-26

The droplet diameter corresponding to a Reynolds number 2-500 was found to be in the range 100:1000 Micrometer. The intermediate law is usually valid for many of the gas liquid and liquid-liquid droplet settling applications encountered in the gas business.

Newton’s Law
Newton’s Law is applicable for a Reynold’s number range of approximately 500 – 200,000, and finds applicability mainly for separation of large droplets or particles from a gas phase, e.g. flare knockout drum sizing.
Vt = 1.74 [g (Dm) (ρd – ρc) / (ρc)]0.5 Eq. 2-27

2.17.2.1: Two phase Separator Sizing
Three main factors should be considered in separator sizing: 1) vapor capacity, 2) liquid capacity, and 3) operability. The vapor capacity will determine the cross-sectional area necessary for gravitational forces to remove the liquid from the vapor. The liquid capacity is typically set by determining the volume required to provide adequate residence time to “de-gas” the liquid or allow immiscible liquid phases to separate. Operability issues include the separator’s ability to deal with solids if present, unsteady flow/liquid slugs, turndown, etc.
Finally, the optimal design will usually result in an aspect ratio that satisfies these requirements in a vessel of reasonable cost. These factors often result in an iterative approach to the calculations.

Separators without Mist Extractors
Separators without mist extractors are not frequently utilized.
The most common application of a vapor-liquid separator that does not use a mist extractor is a flare knockout drum.
Mist extractors are rarely used in flare knockout drums because of the potential for plugging and the serious implications this would have for pressure relief. It is typically a horizontal vessel that utilizes gravity as the sole mechanism for separating the liquid and gas phases. Gas and liquid enter through the inlet nozzle and are slowed to a velocity such that the liquid droplets can fall out of the gas phase. The dry gas passes into the outlet nozzle and the liquid is drained from the lower section of the vessel.
To design a separator without a mist extractor, the minimum size diameter droplet to be removed must be set.
The length of the vessel required can then be calculated by assuming that the time for the gas flow from inlet to outlet is the same as the time for the liquid droplet of diameter Dm to fall from the top of the vessel to the liquid surface. eq. 2-28 then relates the length of the separator to its diameter as a function of this settling velocity (assuming no liquid retention):
L = 4 QA / π Vt Dv Eq. 2-28
If the separator is to be additionally used for liquid storage, this must also be considered in sizing the vessel.
Where
L = Seam to seam Length of separator. ft.
QA = actual gas flow rate, ft3/sec
Vt = Terminal gas velocity, ft./s
Dv = inside diameter of vessel, ft.

yasserkassem · Post by **yasserkassem** » Tue Mar 17, 2020 9:39 pm

Chapter 2 - Part 3

Example 2-7
A horizontal gravity separator (without mist extractor) is required to handle 60 MMscfd of 0.75 specific gravity gas (MW = 21.72) at a pressure of 500 psig and a temperature of 100°F. Compressibility is 0.9, viscosity is 0.012 cp, and liquid specific gravity is 0.5. It is desired to remove all entrainment greater than 150 microns in diameter. No liquid surge is required.
Solution:
Gas density “continues phase”, from eq. 1-16
ρc (gas) = 0.093 ((MW)P)/TZ lb/ft3

= 0.093 x 21.72 x 514.7 / (560 x 0.90) = 2.06 lb/ft3

Liquid droplet density, ρd = 0.5 (62.4) = 31.2 lb/ft3
(ρd – ρc) = 31.2 – 2.06 = 29.14
Droplet diam. (ft.) = 150 x 3.28 x 10-6 = 0.000492
viscosity = 0.012 cp
Volume of gas at operating conditions = V = nzRT/P
Remember (n = ft3 /379)
60 x 106 x 0.9 x 10.7 x 560 / ( 514.7 x 397) = 1.66 x 106 ft3

Volumetric flow rate QA = 1.66 x 106 / ( 24 x 3600) = 19.2 ft3 / sec

Gas velocity = Vol. flow rate QA / area

In case of 3.5 ft diameter, and no liquid volume inside,
Area = 3.14 x (3.5)2 /4 = 9.61 ft2
Velocity = Volumetric flow rate/ area
=19.2 / 9.61 = 2 ft/sec
From eq. 2-23
Re = 1488 DmVt ρc / µ

Re = 1488 x 0.000492 x 2 x 2.06 /0.012
Re = 253
The intermediate law will be used
Vt = [3.49 g0.71 (Dp)1.14 (ρd – ρc)0.71 ] / [(ρc)0.29 µ0.43]

= 3.49 x 11.76 x 1.694 x10-4 x 10.96 / (1.23 x 0.149 )
Vt = 0.42 ft/s
Vessel length = (assume diameter 3.5 ft.)
Using eq. 2-28
L = 4 QA / π Vt Dv
L = 4 x 19.2 / ( 3.14 x 0.46 x 3.5 ) = 15.2 ft

Slenderness ratio = 15.2/3.5 = 4.3
Recalculate 4 ft diameter “48 in.”
Slenderness ratio = 13.3/4 = 3.3
Other D and L combination can be obtained, as in table 2-12,

Table. 2-12. Example 2-7 results.
The user may select a standard vertical separator from table 2-10.

An alternative method is to get Drag coefficient from eq.2-24, extract CD from fig.2-54, and then apply in eq. 2-22, as follows:
Using eq. 2-24
CD (Re)2 = (0.95) (108) ρc (Dm)3 (ρd – ρc )/ µ2
CD (Re)2 = (0.95) (108) (2.07) (0.119 x 10-9) (29.13 )/ 0.000144
CD (Re)2 = 4733
From Fig. 2-54 CD = 1.4
Apply in eq. 2-22.Terminal velocity
Vt = [4gDp (ρd – ρc ) / (3 ρc CD)]0.5

= (4 x 32.2 x 0.000492 x 29.13) / (3 x 2.07 x 1.4) ]0.5
Vt = 0.46 ft/s (Approximately the same result obtained by using previous methods “equations”).
Then continue calculation-using equation 2-28.

Another alternative calculation method as follows:
Max droplet distance = 3.5 ft.
velocity 0.46 ft/s
Time to reach end point = 3.5/0.46 = 7.6 seconds
For droplet settling to occur, gas retention time must be higher than 7.6 sec
Assume 10 ft. length
Volume of vessel will be
10 x (3.5)2 x 3.14/4 = 96 ft3
Retention time (seconds) = volume / flow rate
Retention time (seconds) = 96 / 19.2 = 5 seconds (not enough).
Assume 16 ft. length
Volume of vessel will be
16 x (3.5)2 x 3.14/4 = 154 ft3
Retention time = 154 / 19.2 = 8 seconds (enough).
Select from table 2-10 a standard separator size with diameter 3.5 ft, and length enough for more than 7.6 seconds.
In case of 15.2 ft. length, and 3.5 ft. diameter, retention time will be 7.6 seconds.
Assume different values of (D) , 4, 4.5 , 5 ft. and recalculate to attain more alternatives.

Example 2-8
What size vertical separator without a mist extractor for previous example.
Solution
Gas flow area A = QA/Vt
= 19.2/0.46 = 42 ft2
Area = 42 = D2 x 3.14 /4
D = 7.3 ft = 88 in.
Use 90 in as minimum diameter.

B- Separators with Mist Extractors
Of the four major components of a separator that were discussed in a previous section, the mist extractor has the most impact on separated gas quality with respect to carried over liquid content. The sizing equations and parameters provided in the mist extraction section size the mist extractor itself, not the actual separation vessel. The gas capacities of the various types of mist extractors is generally inversely related to the amount of entrained liquid that the mist extractor is required to remove.

Vertical Separators with Mist Extractors
Gas handling capacity of conventional vertical separators that employ mist extractors has normally been calculated from the Souders and Brown equation, Eq 2-1, using “experience-based” K factors. Typical K values for vertical and horizontal separators from API 12J are presented in table 2-13.

Table 2- 13, K factor for determining maximum allowable superficial velocity.

In qualitative terms, the ranges of K given above may be taken to reflect difficulty of the separation conditions, i.e. from non-ideal/difficult to ideal/easy. As indicated in table 2-13, K is also a function of vessel height. This reflects the fact that a certain minimum distance is required to establish a relatively uniform velocity profile before the gas reaches the mist extractor.

Horizontal Separators with Mist Extractors
Eq. 2-1 can also be used for calculating the gas capacity of horizontal separators.
In calculating the gas capacity of horizontal separators, the cross-sectional area of that portion of the vessel occupied by liquid (at maximum level) is subtracted from the total vessel cross-sectional area. Typical horizontal separator designs will have the normal liquid level at the half-full point. Values of K for horizontal separators from API 12J are given in table 2-12.
In practice, K = 0.5 ft/sec is normally used as an upper limit for horizontal separators equipped with wire-mesh mist extractors. Separators equipped with vane type or cyclonic mist extractors may utilize higher K values than those for mesh pads.
Example 2-9
Gas flow rate 25 MMscfd - Oil Flow rate 3000 BPD
Operating pressure 800 psig - Operating temperature 80 0F
Gas density 3.4 lb/ft3 -- Flowing oil density 51.5 lb/ft3
Compressibility factor z = 0.92 - Separator type Vertical, Two-phase.
Solution:
Assume 10 ft. seam to seam length, 30% liquid full, use K value of 0.3 (table 2-13), and use equation 2-1.
The maximum allowable superficial velocity of the gas is:
Vt = K [(ρl - ρg ) / ρg]0.5
Vt = 0.3 [(51.5 – 3.4 ) / 3.4]0.5 = 1.128 ft/sec
Actual volume flow rate of gas, (V = nzRT/P)
25 x 1006 x 0.92 x 10.73 x 540 / (379 x 814.7) = 431603 ft3/day.
= 5.0 ft3/sec
Minimum gas flow area = 5 / 1.128 = 4.43 ft2
Gas flow area = π D2 /4 = 3.14 x D2 /4 = 4.43 (ft2)
Minimum separator ID, D = 2.38 ft
Minimum diameter, in. (d) = 29 in.
Use 30 in, vessel diameter
Assume 1 minute retention time for liquid, from table 2-8.
Liquid volume = π D2 h /4
(h = 3 ft. “30% of 10 ft.”)
= 3.14 x 302 x 3 / (144 x 4) = 14.7 ft3
= 14.7 x 0.178 = 2.62 bbl
The liquid capacity of the separator from eq. 2-5
2.62 = (W x 1)/1440 w = 3773 bpd
So, liquid capacity is satisfactory for 30 in. and 10 ft.

Calculation tips
In separators design, the residence time of gas, must be higher than the time required for a certain droplet diameter to fall. In addition, the residence time of liquids must be higher than the time required for a certain volume of gas bubbles to evolve. In three-phase separation, The liquid residence time must be higher than the time required for a certain size of water droplets to fall from oil phase to water phase, and for oil droplets trapped in water to evolve to the oil phase.
To check the above condition for existing vessel, or for selected dimensions, all you have to do is as follows:
a- Calculate the distance that droplets or bubble has to travel. (for example in horizontal separator: =d/2 for liquid in gas and vapor in liquid, and less for water in oil and oil in water)
b- Calculate the velocity of falling or evolving (you may use intermediate law directly)
c- Calculate time required for falling or evolving (distance/Velocity)
c- Calculate residence time of liquid, and for gas at operating conditions.
d- Residence time must be higher than droplet falling time, or bubble or oil in water evolving time.

yasserkassem · Post by **yasserkassem** » Fri Mar 20, 2020 3:03 pm

Chapter 3

----
Chapter 3 107
Three-phase Oil and Gas Separation 107
3.1: Introduction 107
3.2: three phase separation equipment’s 108
3.2.1: Horizontal Separators 108
3.2.2: Free-Water Knockout 111
3.2.3: Horizontal Three-Phase Separator with a Liquid “Boot” 111
3.2.4: Vertical Separators 112
3.2.5: Selection Considerations 114
3.3: Internal Vessel components 115
3.3.1: Coalescing Plates 117
3.4: Operating Problems 118
3.4.1: Emulsions 118
3.5: Three-Phase Separator Design Theory 118
3.5.1: Gas Separation 118
3.5.2: Oil–Water Settling 118
3.5.3: Water Droplet Size in Oil 118
3.5.4: Oil Droplet Size in Water 119
3.5.5: Retention Time 119
3.6: Separator Design (first method) 121
3.6.1: Horizontal Three-phase Separator Sizing—Half-Full 121
3.6.1.2: Retention Time Constraint 121
3.6.1.3: Settling Water Droplets from Oil Phase 122
3.6.1.4: Separating Oil Droplets from Water Phase 123
3.6.2: Vertical Separators’ Sizing 124
3.6.2.1: Gas Capacity Constraint 125
3.6.2.3: Settling Oil from Water Phase Constraint 125
3.7: Separator Design (second method) 131
---------------------

Chapter 3

Three-phase Oil and Gas Separation

3.1: Introduction
When oil and water are mixed with some intensity and then allowed to settle, a layer of relatively clean free water will appear at the bottom.
The growth of this water layer with time will follow a curve as shown in Figure 3-1.

Fig. 3-1. Growth of water layer with time.

After a period of time, ranging anywhere from 3 minutes to 30 minutes, the change in the water height will be negligible. The water fraction, obtained from gravity settling, is called “free water.” It is normally beneficial to separate the free water before attempting to treat the remaining oil and emulsion layers.
“Three-phase separator” and “free-water knockout” are terms used to describe pressure vessels that are designed to separate and remove the free water from a mixture of crude oil and water.
Because flow normally enters these vessels directly from either (1) a producing well or (2) a separator operating at a higher pressure, the vessel must be designed to separate the gas that flashes from the liquid as well as separate the oil and water.
The term “three-phase separator” is normally used when there is a large amount of gas to be separated from the liquid, and the dimensions of the vessel are determined by the gas capacity equations discussed in previous chapter. While, “Free-water knockout” is generally used when the amount of gas is small relative to the amount of oil and water, and the dimensions of the vessel are determined by the oil–water separation equations will be discussed in this chapter. No matter what name is given to the vessel, any vessel that is designed to separate gas in addition of separation of two immiscible liquid phases will employ the concepts described in this chapter, and we will call such a vessel a “three-phase separator.”
The basic design aspects of three-phase separation are identical to those discussed for two-phase separation in Chapter 2. The only additions are that more concern is placed on liquid-liquid settling rates and that some means of removing the free water must be added. Liquid-liquid settling rates will be discussed later in this chapter.
3.2: three phase separation equipment’s
Three-phase separators are designed as either horizontal or vertical pressure vessels.
3.2.1: Horizontal Separators
Figures 3-2 and 3-3 are schematics of horizontal three-phase separator. The fluid enters the separator and hits an inlet diverter. This sudden change in momentum does the initial gross separation of liquid and vapor as discussed in Chapter 2. In most designs the inlet diverter contains a down-comer that directs the liquid flow below the oil–water interface.
The down-comer forces the inlet mixture of oil and water to mix with the water continuous phase in the bottom of the vessel and rise through the oil–water interface. This process is called “water washing,” and it promotes the coalescence of water droplets, which are entrained in the oil continuous phase. Figure 3-4 illustrates the principles of “water washing.” The inlet diverter assures that little gas is carried with the liquid, and the water wash assures that the liquid does not fall on top of the gas–oil or oil–water interface, mixing the liquid retained in the vessel and making control of the oil–water interface difficult.

Fig.3-2. Simple horizontal three-phase separator sketch.

Fig. 3-3. Horizontal three-phase separator with interface level control and weir.

Figure 3-4. Inlet diverter illustrating the principles of “water washing”

The liquid collecting section of the vessel provides sufficient time so that the oil and emulsion form a layer or “oil pad” on top of the free water. The free water settles to the bottom. Figure 3-3 is a horizontal three-phase separator with an interface level controller and weir. The weir maintains the oil level, and the level controller maintains the water level. The oil is skimmed over the weir. A level controller that operates the oil dump valve controls the level of the oil downstream of the weir.
The produced water flows from a nozzle in the vessel located upstream of the oil weir. An interface level controller senses the height of the oil–water interface. The controller sends a signal to the water dump valve, thus allowing the correct amount of water to leave the vessel so that the oil–water interface is maintained at the design height.
The gas flows horizontally and out through a mist extractor to a pressure control valve that maintains constant vessel pressure. The level of the gas–oil interface can vary from 50% to 75% of the diameter depending on the relative importance of liquid–gas separation. The most common configuration is half-full, and this is used for the design equations in this section. Similar equations can be developed for other interface levels. Figure 3-5 shows an alternate configuration known as a “bucket and weir” design. Figure 3-6 is a cutaway view of a horizontal three-phase separator with a bucket and weir. This design eliminates the need for a liquid interface controller.
Both the oil and water flow over weirs where level control is accomplished by a simple displacer float. The oil overflows the oil weir into an oil bucket where its level is controlled by a level controller that operates the oil dump valve. The water flows under the oil bucket and then over a water weir. The level downstream of this weir is controlled by a level controller that operates the water dump valve.
As shown in Figures 3-5 and 3-6, the back of the oil bucket is higher than the front of the bucket. This differential height configuration assures oil will not flow over the back of the bucket and out with the water should (if) the bucket become flooded.
The height of the oil weir controls the liquid level in the vessel. The difference in height of the oil and water weirs controls the thickness of the oil pad due to specific gravity differences. It is critical to the operation of the vessel that the water weir height is sufficiently below the back oil weir height so that the oil pad thickness provides sufficient oil retention time. If the water weir is too low and the difference in specific gravity is not as great as anticipated, then the oil pad could grow in thickness to a point where oil will be swept under the oil box and out the water outlet. Normally, either the oil or the water weir is made adjustable so that changes in oil or water specific gravities or flow rates can be accommodated.

Figure 3-5. Horizontal three-phase separator with a “bucket and weir.”

Figure 3-6. Horizontal three-phase separator with a “bucket and weir.”

To obtain a desired oil pad height, the water weir should be set a distance below the oil weir. This distance is calculated by using Eq. (3-1), which is developed by equating the static heads at point “A.”

Δh = ho [ 1- (ρo/ρw) ] Eq. 3-1
where
Δh = distance below the oil weir, in,
ho = desired oil pad height, in,
ρo = oil density, lb/ft3,
ρw = water density, lb/ft3.
This equation neglects the height of the oil and water flowing over the weir and presents a view of the levels when there is no inflow. A large inflow of oil will cause the top of the oil pad to rise; the oil pad will thus get thicker, and the oil bucket must be deep enough so that oil does not flow under it. Similarly, a large inflow of water will cause the level of water flowing over the water weir to rise, and there will be a large flow of oil from the oil pad over the oil weir until a new hw is established (Fig.3-7).
These dynamic effects can be minimized by making the weirs as long as possible.
Three-phase separators with a bucket and weir design are most effective with high water-to-oil flow rates and/or small density differences.
Interface control design has the advantage of being easily adjustable to handle unexpected changes in oil or water specific gravity or flow rates.
Interface control should be considered for applications with high oil flow rates and/or large density differences. However, in heavy oil applications or where large amounts of emulsion or paraffin are anticipated, it may be difficult to sense interface level. In such a case bucket and weir control is recommended.

Fig.3-7. Determination of oil pad height.

3.2.2: Free-Water Knockout
Free Water Knockout is a three-phase separator which is used to remove free water held in the vessel and separate brine from crude oil. The term (FWKO) is reserved for a vessel that processes an inlet liquid stream with little entrained gas and makes no attempt to separate the gas from the oil.
Figure 3-8 illustrates a horizontal FWKO. Figure 3-9 illustrates a vertical FWKO.
The major difference between a conventional three-phase separator and an FWKO is that in the latter there are only two fluid outlets; one for oil and very small amounts of gas and the second for the water. FWKOs are usually operated as packed vessels. Water outflow is usually controlled with an interface level control. The design of an FWKO is the same as that of a three-phase separator. Since there is very little gas, the liquid capacity constraint always dictates the size.

3.2.3: Horizontal Three-Phase Separator with a Liquid “Boot”
Figures 3-10 and 3-11, show horizontal three-phase separator with a water “boot” on the bottom of the vessel barrel. The “boot” collects small amounts of water that settle out in the liquid collection section and travel to the outlet end of the vessel. These vessels are a special case of three-phase separators. In this case, the flow rate of both oil and water can provide enough retention time for separation of oil and water, with a little possibility that emulsion or crude oil escape through the water drain.

Figure 3-8. Horizontal FWKO.

Figure 3-9. Vertical FWKO.

3.2.4: Vertical Separators
In vertical three-phase separator figures 3-12 and 3-13, the flow enters the vessel through the side as in the horizontal separator. The inlet diverter separates the bulk of the gas. A down-comer is required to route the liquid through the oil–gas interface so as not to disturb the oil skimming action taking place. A chimney is needed to equalize gas pressure between the lower section and the gas section.
The spreader, or down-comer, outlet is located just below the oil–water interface, thus “water washing” the incoming stream. From this point as the oil rises, any free water trapped within the oil phase separates out.
The water droplets flow countercurrent to the oil. Similarly, the water flows downward and oil droplets trapped in the water phase tend to rise countercurrent to the water flow.

Figure 3-10. Horizontal three-phase separator with a “water boot.”

Figure 3-11. Horizontal three-phase separator with a “water boot.”

Figures 3-14 and 3-15 are views of vertical three-phase separators without water washing and with interface control.
Three different methods of control that are often used on vertical separators, fig 3-16.
The first uses a regular displacer float is used to control the gas–oil interface and regulate a control valve dumping oil from the oil section. An interface float is used to control the oil–water interface and regulate a water outlet control valve. Because no internal baffling or weirs are used, this system is the easiest to fabricate and handles sand and solids production best.
The second method uses a weir to control the gas–oil interface level at a constant position. This results in a better separation of water from the oil as all the oil must rise to the height of the oil weir before exiting the vessel. Its disadvantages are that the oil box takes up vessel volume and costs money to fabricate. In addition, sediment and solids could collect in the oil box and be difficult to drain, and a separate low-level shut-down may be required to guard against the oil dump valve’s failing to open.
The third method uses two weirs, which eliminates the need for an interface float. Interface level is controlled by the height of the external water weir relative to the oil weir or outlet height. This is similar to the bucket and weir design of horizontal separators. The advantage of this system is that it eliminates the interface level control. The disadvantage is that it requires additional external piping and space. In cold climates the water leg is sometimes installed internal to the vessel so that the vessel insulation will prevent it from freezing.

Figure 3-12. Vertical three-phase separator with interface level control. Figure 3-13. Vertical three-phase separator with interface level control.

3.2.5: Selection Considerations
The geometry and physical and operating characteristics give each separator type advantages and disadvantages.
Gravity separation is more efficient in horizontal vessels than in vertical vessels.
In the gravity settling section of a horizontal vessel, the settling velocity and flow velocity are perpendicular rather than countercurrent in a vertical vessel.
Horizontal separators have greater interface areas, which enhances phase equilibrium. This is especially true if foam or emulsion collect at the gas–oil interface.
Thus, from a process perspective, horizontal vessels are preferred. However, they do have several drawbacks, which could lead to a preference for a vertical vessel in certain situations:
Horizontal separators are not as good as vertical separators in handling solids. The liquid dump valve of a vertical separator can be placed at the center of the bottom head so that solids will not build up in the separator, but continue to the next vessel in the process. As an alternative, a drain could be placed at this location so that solids could be disposed of periodically while liquid leaves the vessel at a slightly higher elevation.
In a horizontal vessel, it is necessary to place several drains along the length of the vessel. Since the solids will have an angle of repose of 450 to 600, the drains must be spaced at very close intervals [usually no farther than 5 ft apart]. Attempts to lengthen the distance between drains, by providing sand jets in the vicinity of each drain to fluidize the solids while the drains are in the operation, are expensive and have been only marginally successful in field operations.
Horizontal vessels require more plan area to perform the same separation as vertical vessels. While this may not be of importance at a land location, it could be very important offshore. If several separators are used, however, this disadvantage may be overcome by stacking horizontal separators on top of each other.
Small-diameter horizontal vessels [3-ft diameter and smaller] have less liquid surge capacity than vertical vessels sized for the same steady-state flow rate. For a given change in liquid surface elevation, there is typically a larger increase in liquid volume for a horizontal separator than for a vertical separator sized for the same flow rate. However, the geometry of a small horizontal vessel causes any high-level shut-down device to be located close to the normal operating level. In very large diameter [greater than 6 ft] horizontal vessels and in vertical vessels, the shut-down could be placed much higher, allowing the level controller and dump valve more time to react to the surge. In addition, surges in horizontal vessels could create internal waves, which could activate a high-level sensor prematurely. Care should be exercised when selecting small-diameter [5 ft] horizontal separators. The level controller and level switch elevations must be considered. The vessel must have a sufficiently large diameter so that the level switches may be spaced far enough apart, vertically, so as to avoid operating problems. This is important if surges in the flow of slugs of liquids are expected to enter the separator.

It should be pointed out that vertical vessels have some drawbacks that are not process related and that must be considered when making a selection. For example, the relief valve and some of the controls may be difficult to service without special ladders and platforms. The vessel may have to be removed from the skid for trucking due to height restrictions.

In summary, horizontal vessels are most economical for normal oil–water separation, particularly where there may be problems with emulsions, foam, or high gas–liquid ratios. Vertical vessels work most effectively in low gas–oil ratio (GOR) applications and where solids production is anticipated.

3.3: Internal Vessel components
Vessel internals common to both two-phase and three-phase separators, such as inlet diverters, wave breakers, de-foaming plates, vortex breakers, stilling wells, sand jets and drains, and mist extractors, are covered in previous Chapter and will not be repeated here. Additional internals that aid in the separation of oil and water are presented in this section.

Figure 3-14. Vertical three-phase separator without water washing and with vane mist extractor.

Figure 3-15. Vertical three-phase separator without water washing and with wire-mesh mist extractor.

Figure 3-16. Liquid level control for three-phase vertical separators.

3.3.1: Coalescing Plates
It is possible to use various plate or pipe coalescer designs to aid in the coalescing of oil droplets in the water and water droplets in the oil. The installation of coalescing plates in the liquid section will cause the size of the water droplets entrained in the oil phase to increase, making gravity settling of these drops to the oil–water interface easier. Thus, the use of coalescing plates (Figure 3-17), will often lead to the ability to handle a given flow rate in a smaller vessel.
However, because of the potential for plugging with sand, paraffin, or corrosion products, the use of coalescing plates should be discouraged, except for instances where the savings in vessel size and weight are large enough to justify the potential increase in operating costs and decrease in availability.

Figure 3-17. Horizontal three-phase separator fitted with coalescing plates.

3.4: Operating Problems
3.4.1: Emulsions
Three-phase separators may experience the same operating problems as two-phase separators. In addition, three-phase separators may develop problems with emulsions which can be particularly troublesome in the operation of three-phase separators. Over a period of time an accumulation of emulsified materials and/or other impurities may form at the interface of the water and oil phases. In addition to adverse effects on the liquid level control, this accumulation will also decrease the effective oil or water retention time in the separator, with a resultant decrease in water–oil separation efficiency. Addition of chemicals and/or heat often minimizes this difficulty.
Frequently, it is possible to appreciably lower the settling time necessary for water–oil separation by either the application of heat in the liquid section of the separator or the addition of de-emulsifying chemicals. The treating of emulsions is discussed in more detail in next chapter.
3.5: Three-Phase Separator Design Theory
3.5.1: Gas Separation
The concepts and equations pertaining to two-phase separation described in previous chapter are equally valid for three-phase separation.

3.5.2: Oil–Water Settling
It can be shown that flow around settling oil drops in water or water drops in oil is laminar and thus Stokes’ law governs. The terminal drop velocity is

Vt = 1.78 x 10-6 (ρd – ρc ) d2m / µ Eq. 3-2

ρd =density of droplet fluid, lb/ft3, (= 62.4×SG)
ρc = density of the continuous phase (Medium), lb/ft3, (= 62.4×SG)
In case of separation of water droplet in oil, continuous phase is oil and droplet is water, while, In case of separation of oil droplet in water phase, Vt value will be negative, since movement of oil droplet will be upward.
Vt =terminal (settling velocity) of the droplet, ft/s,
dm =droplet diameter, microns,
µ =viscosity of the continuous phase, cp.

3.5.3: Water Droplet Size in Oil
It is difficult to predict the water droplet size that must be settled out of the oil phase to coincide with the rather loose definition of “free oil.”
Unless laboratory or nearby field data are available, good results have been obtained by sizing the oil pad such that water droplets 500 microns and larger settle out. As shown in Figure 3-18, if this criterion is met, the emulsion to be treated by downstream equipment should contain less than 5% to 10% water. In heavy crude oil systems, it is sometimes necessary to design for 1,000-micron water droplets to settle. In such cases the emulsion may contain as much as 20% to 30% water.

3.5.4: Oil Droplet Size in Water
From Eq. (3-2) it can be seen that the separation of oil droplets from the water is easier than the separation of water droplets from the oil. The oil’s viscosity is on the order of 5 to 20 times that of water. Thus, the terminal settling velocity of an oil droplet in water is much larger than that of a water droplet in oil. The primary purpose of three phase separation is to prepare the oil for further treating. Field experience indicates that oil content in the produced water from a three-phase separator, sized for water removal from oil, can be expected to be between a few hundred and 2,000 mg/l. This water will require further treating prior to disposal. Sizing for oil droplet removal from the water phase does not appear to be a meaningful criterion.
Occasionally, the viscosity of the water phase may be as high as, or higher than, the liquid hydrocarbon phase viscosity. For example, large glycol dehydration systems usually have a three-phase flash separator.
The viscosity of the glycol/water phase may be rather high. In cases like this, the settling equation should be applied to removing oil droplets of approximately 200 microns from the water phase.
If the retention time of the water phase is significantly less than the oil phase, then the vessel size should be checked for oil removal from the water. For these reasons, the equations are provided so the water phase may be checked. However, the separation of oil from the water phase rarely governs the vessel size and may be ignored for most cases.

3.5.5: Retention Time
A certain amount of oil storage is required to assure that the oil reaches equilibrium and that flashed gas is liberated. An additional amount of storage is required to assure that the free water has time to coalesce into droplet sizes sufficient to fall in accordance with Eq. 3-2.
It is common to use retention times ranging from 3 minutes to 30 minutes depending upon laboratory or field data. If this information is not available, the guidelines presented in Table 3-1, can be used. Generally, the retention time must be increased as the oil gravity or viscosity increases.
Similarly, a certain amount of water storage is required to assure that most of the large droplets of oil entrained in the water have sufficient time to coalesce and rise to the oil–water interface. It is common to use retention times for the water phase ranging from 3 minutes to 30 minutes depending upon laboratory or field data. If this information is not available, a water retention time of 10 minutes is recommended for design.
The retention time for both the maximum oil rate and the maximum water rate should be calculated, unless laboratory data indicate that it is unnecessary to take this conservative design approach.

Table 3-1. API 12J, recommended oil retention time.

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Figure 3-18. Example water droplet size distribution. Size distribution varies widely for different process conditions and crude and water properties.

3.6: Separator Design (first method)
The guidelines presented here can be used for initial sizing of a horizontal three-phase separator 50% full of liquid.

3.6.1: Horizontal Three-phase Separator Sizing—Half-Full
For sizing a horizontal three-phase separator it is necessary to specify a vessel diameter and a seam-to-seam vessel length. The gas capacity and retention time considerations establish certain acceptable combinations of diameter and length. The need to settle 500-micron water droplets from the oil and 200-micron oil droplets from the water establishes a maximum diameter corresponding to the given liquid retention time.

3.6.1.1: Gas Capacity Constraint
The principles of liquid droplets settling through a gas, which were derived in previous chapter, can be used to develop an equation to size a separator for a gas flow rate. By setting the gas retention time equal to the time required for a drop to settle to the liquid interface, the following equations may be derived:
dLeff = 420 (TZQ/P) [CD ρ1/(ρd- ρc) dm ]0.5 Eq. 3-3
where
d =vessel internal diameter, in,
Leff =effective length of the vessel where separation occurs, ft,
T =operating temperature, 0R,
Qg =gas flow rate, MMscfd,
P =operating pressure, psia,
Z =gas compressibility,
CD =drag coefficient,
dm =liquid droplet to be separated, micron,
ρc=density of continuous phase “medium” (gas in this case), lb/ft3,
ρd=density of droplet, lb/ft3.

3.6.1.2: Retention Time Constraint
Liquid retention time constraints can be used to develop the following equation, which may be used to determine acceptable combinations of d and Leff .

d2Leff = 1.42 [(Qw)(tr)w + (Qo)(tr)o] Eq. 3-4

where
Qw = water flow rate, BPD
(tr)w = water retention time, min,
Qo = oil flow rate, BPD,
(tr)o = oil retention time, min,

3.6.1.3: Settling Water Droplets from Oil Phase
The velocity of water droplets settling through oil can be calculated using Stokes’ law. From this velocity and the specified oil phase retention time, the distance that a water droplet can settle may be determined. This settling distance establishes a maximum oil pad thickness given by the following formula:
(ho)max = 0.00128 (tr)o (ΔSG ) d2m / µ Eq. 3-5
ΔSG = difference in sp.gr Kg/l.
This is the maximum thickness the oil pad can be and still allow the water droplets to settle out in time (tr)o. For dm = 500 microns, the following equation may be used.

(ho)max = 320 (tr)o (ΔSG) / µ Eq. 3-6

For a given oil retention time [(tr)o] and a given water retention time [(tr)w], the maximum oil pad thickness constraint establishes a maximum diameter in accordance with the following procedure:
1. Compute (ho)max. Use 500-micron droplet if no other information is available.
2. Calculate the fraction of the vessel cross-sectional area occupied by the water phase. This is by
Aw/A = 0.5 Qw (tr)w / [ (tr)o Qo + (tr)w Qw ] Eq. 3-7

Figure 3-19. Coefficient “β ” for a cylinder half filled with liquid.

3. From Figure 3-19, determine the coefficient β.
4. Calculate dmax from
dmax = (ho)max/β Eq. 3-8
where
β = ho/d
Any combination of d and Leff that satisfies all three of Eqs. (3-3), (3-4), and (3-5) will meet the necessary criteria.

3.6.1.4: Separating Oil Droplets from Water Phase
Oil droplets in the water phase rise at a terminal velocity defined by Stokes’ law. As with water droplets in oil, the velocity and retention time may be used to determine a maximum vessel diameter from Eqn. 3-4.
It is rare that the maximum diameter determined from a 200-micron oil droplet rising through the water phase is larger than a 500-micron water droplet falling through the oil phase. Therefore, the maximum diameter determined from a 500-micron water droplet settling through the oil phase normally governs the vessel design. For dm = 200 microns, the following equations may be used:

(hw)max = 51.2 (tr)w (ΔSG) / µw Eq. 3-9

ΔSG = difference in sp.gr Kg/l.
The maximum diameter may be found from the following equation:

dmax = (hw)max / β Eq. 3-10

3.6.1.5: Seam-to-Seam Length
The effective length may be calculated from Eq. 3-4. From this, a vessel seam-to-seam length may be estimated. The actual required seam-to-seam length is dependent on the physical design of the vessel.
For vessels sized based on gas capacity, some portion of the vessel length is required to distribute the flow evenly near the inlet diverter.
Another portion of the vessel length is required for the mist extractor.
The length of the vessel between the inlet and the mist extractor with evenly distributed flow is the Leff calculated from Eq. 3-3. As a vessel’s diameter increases, more length is required to evenly distribute the gas flow. However, no matter how small the diameter may be, a portion of the length is still required for the mist extractor and flow distribution. Based on these concepts coupled with field experience, the seam-to-seam length of a vessel may be estimated as the larger of the following:
Lss = 4/3 Leff Eq. 3-11
Lss = Leff + d/12 Eq. 3-12

For vessels sized on a liquid capacity basis, some portion of the vessel length is required for inlet diverter flow distribution and liquid outlet. The seam-to-seam length is estimated as follow:
Lss = 4/3 Leff Eq. 3-13

3.6.1.6: Slenderness Ratio
For each vessel design, a combination of Leff and d exists that will minimize the cost of the vessel. In general, the smaller the diameter of a vessel, the less it will cost. However, decreasing the diameter increases the fluid velocities and turbulence. As a vessel diameter decreases, the likelihood of the gas re-entraining liquids or destruction of the oil/water interface increases. Experience indicates that the ratio of the seam-to seam length divided by the outside diameter should be between 3 and 5. This ratio is referred to as the “slenderness ratio” (SR) of the vessel.

3.6.1.7: Procedure for Sizing Three-Phase Horizontal Separators—Half-Full
1. The first step in sizing a horizontal separator is to establish the design basis. This includes specifying the maximum and minimum flow rates, operating pressure and temperature, droplet size to be removed, etc.
2. Select a (tr)o and a (tr)w.

3. Calculate (ho)max. from eq.3-5, Use a 500-micron droplet if no other information is available.
(ho)max= 0.00128 (tr)o (ΔSG ) d2m / µ Eq.3-5
For 500 micron use eq. 3-6.
(ho)max = 320 (tr)o (ΔSG) / µ Eq. 3-6

4. Calculate Aw/A: use eq. 3-7.
Aw/A = 0.5 Qw (tr)w / [ (tr)o Qo + (tr)w Qw ] Eq.3-7

5. Determine β from curve.

6. Calculate dmax: use eq. 3-8
dmax = (ho)max/β Eq. 3-8

7. Calculate combinations of d, Leff for d less than dmax that satisfy the gas capacity constraint. Use 100-micron droplet if no other information is available.use eq. 3-3.
dLeff = 420 (TZQ/P) [CD ρc/(ρd- ρc) dm ]0.5 Eq. 3-3

8. Calculate combinations of d, Leff for d less than dmax that satisfy the oil and water retention time constraints. Use eq. 3-4
d2Leff = 1.42 [(Qw)(tr)w + (Qo)(tr)o] Eq. 3-4

9. Estimate seam-to-seam length. Use eqs. 3-11 or 3-12.
Lss = Leff + d/12 (gas capacity) Eq. 3-11

Lss = 4/3 Leff (liquid capacity) Eq. 3-12

10. Select a reasonable diameter and length. Slenderness ratios (12Lss/d) on the order of 3 to 5 are common.
11. When making a final selection, it is always more economical to select a standard vessel size. Table 2-9 in previous chapter, represents API Spec. 12J for horizontal separator standard sizes.

3.6.2: Vertical Separators’ Sizing
As with vertical two-phase separators, a minimum diameter must be maintained to allow liquid droplets to separate from the vertically moving gas. The vessel must also have a large enough diameter to allow water droplets to settle in the upward-flowing oil phase and to allow oil droplets to rise in the downward-moving water phase. The liquid retention time requirement specifies a combination of diameter and liquid volume height.
Any diameter greater than the minimum required for gas capacity and for liquid separation can be chosen.

3.6.2.1: Gas Capacity Constraint
By setting the gas velocity equal to the terminal settling velocity of a droplet, the following may be derived:

d2 = 5040 (TZQg/P) [CD ρc/(ρd- ρc) dm ]0.5 Eq. 3-14
where
d =vessel internal diameter, in.,
T =operating temperature, 0R,
Qg =gas flow rate, MMscfd,
P =operating pressure, psia,
Z =gas compressibility,
CD =drag coefficient,
dm =liquid droplet to be separated, micron,
ρc=density of continues phase “gas in this case”, lb/ft3,
ρd=density of droplet, lb/ft3.

For 100-micron droplet removal, Eqs. 3-14 is reduced to the following:

d2 = 504 (TZQg/P) [CD ρc/(ρd- ρc) ]0.5 Eq. 3-15

3.6.2.2: Settling Water Droplets from Oil Phase Constraint
The requirement for settling water droplets from the oil requires that the following equation must be satisfied:

d2 = 6690 Qo µ/ (ΔSG) d2m Eq. 3-16

ΔSG = difference in sp.gr Kg/l.
for 500 micron water droplet eq. 3-16 can be

d2 = 0.0267 Qo µ/ (ΔSG) Eq. 3-17

3.6.2.3: Settling Oil from Water Phase Constraint

The requirement for separating oil from water requires that the following equation must be satisfied:
d2 = 6690 Qw µ/ (ΔSG) d2m Eq. 3-18

ΔSG = difference in sp.gr Kg/l.
for 200 micron oil droplet eq. 3-18 can be

d2 = 0.167 Qw µ/ (ΔSG) Eq. 3-19

3.6.2.4: Retention time constraint
ho +hw = [(tr)oQo + (tr)w Qw] / 0.12d2 Eq. 3-20
where
ho = height of oil pad, in.,
hw = height from water outlet to interface, in.

3.6.2.5: Seam-to-Seam Length
As with horizontal three-phase separators, the specific design of the vessel internals will affect the seam-to-seam length. The seam-to-seam length (Lss) of vertical vessels may be estimated based on the diameter and liquid height. As shown in Figure 3-20, allowance must be made for the gravity settling (gas separation) section, inlet diverter, mist extractor, and any space below the water outlet. For screening purposes, the Lss values are given from one of eqs. 3-21 and 3-22.
Lss = (h0+ hw+76)/12 for diameters <=36 in. Eq. 3-21
Lss = (h0+ hw+d+40)/12 for diameters >36 in. Eq. 3-22

3.6.2.6: Slenderness Ratio
As with horizontal three-phase separators, the larger the slenderness ratio, the less expensive the vessel. In vertical separators whose sizing is liquid dominated, it is common to choose slenderness ratios no greater than 4 to keep the height of the liquid collection section to a reasonable level. Choices between1.5 to 3 are common, although height restrictions may force the choice of a lower slenderness ratio.

3.6.2.7: Procedure for Sizing Three-Phase Vertical Separators
1. The first step in sizing a vertical separator is to establish the design basis. This includes specifying the maximum and minimum flow rates, operating pressure and temperature, droplet size to be removed, etc.
2. Equation 3-14 may be used to calculate the minimum diameter for a liquid droplet to fall through the gas phase.
Use Eq. 3-15 for 100-micron droplets if no other information is available.
d2 = 5040 (TZQg/P) [CD ρc/(ρd- ρc) dm ]0.5 Eq. 3-14

d2 = 504 (TZQg/P) [CD ρc/(ρd- ρc) ]0.5 Eq. 3-15

3. Equation 3-16 may be used to calculate the minimum diameter for water droplets to fall through the oil phase.
Use Eq. 3-17 for 500-micron droplets if no other information is available.
d2 = 6690 Qo µ/ (ΔSG) d2m Eq. 3-16

d2 = 0.0267 Qo µ/ (ΔSG) Eq. 3-17

4. Equation (3-18) may be used to calculate the minimum diameter for oil droplets to rise through the water phase.
Use Eq. 3-19 for 200-micron droplets if no other information is available.
d2 = 6690 Qw µ/ (ΔSG) d2m Eq. 3-18
d2 = 0.167 Qw µ/ (ΔSG) Eq. 3-19
5. Select the largest of the three diameters calculated in steps 2–4 as the minimum diameter. Any value larger than this minimum may be used for the vessel diameter.
6. For the selected diameter, and assumed values of (tr)o and (tr)w, Eq 3-20 may be used to determine ho+hw.
ho +hw = [(tr)oQo + (tr)w Qw] / 0.12d2 Eq. 3-20

7. From d and ho+hw the seam-to-seam length may be estimated using Eqs. 3-21 and 3-22. The larger value of Lss should be used.

Lss = (h0+ hw+76)/12 for diameters <=36 in. Eq. 3-21
Lss = (h0+ hw+d+40)/12 for diameters >36 in. Eq. 3-22

8. Check the slenderness ratios. Slenderness ratios between 1.5 and 3 are common. The following equations may be used:
SR = 12 Lss/d Eq. 3-23
9. If possible, select a standard-size diameter and seam-to-seam length.

Figure 3-20. Approximate seam-to-seam shell length for a vertical three-phase separator.

Example 3-1: Sizing a vertical three-phase separator.
Given:
Qo = 5000 BOPD,
Qw = 3000 BWPD,
Qg = 5 MMscfd,
P = 100 psia,
Temp. = 900F,
Oil = 30 0API,
(SG)w = 1.07,
Sg = 0.6,
(tr)o = (tr)w = 10 min,
µo = 10 cp,
µw = 1 cp,
ρg = 0.3 lb/ft3,
ρl = 54.7 lb/ft3,
CD = 2.01
Z = 0.99
Droplet removal = 100 microns liquids, 500 microns water, 200 microns oil.

Solution:
1. Calculate difference in specific gravities.
0API (30) = [141.5 /(SG)o]−131.5
(SG)o = 0.876
ΔSG = 1.07 – 0.876 = 0.194
2. Calculate the minimum diameter required to settle a liquid droplet through the gas phase [Eq. (3-14)].
d2 = 5040 (TZQg/P) [CD ρc/(ρd- ρc) dm ]0.5 Eq. 3-14
d2 = 5040 (550x0.99x5/100) [2.01x 0.3/(54.7- 0.3) 100 ]0.5
d = 38 in.
3. Calculate the minimum diameter required for water droplets to settle through the oil phase.
d2 = 6690 Qo µ/ (ΔSG) d2m Eq. 3-16
d2 = 6690 x 5000 x10/ (0.194) (500)2
d = 83 in.

4. Calculate the minimum diameter required for oil droplets to rise through the water phase.
d2 = 6690 Qw µ/ (ΔSG) d2m Eq. 3-16
d2 = 6690 x 3000 x 1 / (0.194) (200)2
d = 50.8 in.

5. Select the largest diameter from steps 2–4 as the minimum inside diameter required.
dmin = 83.0 in.
Choose different standard d ( 84, 90, and 96 in.), and proceed calculation .

6. Calculate ho +hw.
ho +hw = [(tr)oQo + (tr)w Qw] / 0.12d2 Eq. 3-20
For d= 84 in.
ho +hw = [10 x 5000 + 10 x 3000] / 0.12x (84)2
ho +hw = 94.5 in.
For d= 90 in. ho +hw = 82.3 in.
For d= 96 in. ho +hw = 72.3 in.

7. Compute seam-to-seam length (Lss). Select the larger value from Eq. 3-21 and 3-22.
Lss = (h0+ hw+76)/12 for diameters <=36 in. Eq. 3-21
Lss = (h0+ hw+d+40)/12 for diameters >36 in. Eq. 3-22

8. Compute the slenderness ratio.
SR = 12 Lss/d Eq. 3-23

Choices in the range of 1.5 to 3 are common.
Refer to Table 3-2 for results.

Table 3-2, solution of example 1.
9. Make final selection: compute combinations of d and ho+hw for diameters greater than the minimum diameter. See Table 3-2 for results. Select 90 in outside diameter (OD) ×20 ft seam-to-seam length.

Example 3-2: Sizing a horizontal three-phase separator
Given:
Qo = 5000 BOPD, Qw = 3000 BWPD, Qg = 5 MMscfd,
P = 100 psia, Temp. = 900F, Oil = 30 0API,
(SG)w = 1.07,
Sg = 0.6,
(tr)o = (tr)w = 10 min,
µo = 10 cp,
µw = 1 cp,
ρg = 0.3 lb/ft3,
ρl = 54.7 lb/ft3,
CD = 2.01
Z = 0.99
Droplet removal = 100 microns liquids, 500 microns water, 200 microns oil.
Vessel is half-full of liquids

Solution:
1. Calculate difference in specific gravities.
0API (30) = [141.5 /(SG)o]−131.5
(SG)o oil = 0.876
ΔSG = 1.07 – 0.876 = 0.194
2. Calculate maximum oil pad thickness (ho) max. Use 500-micron droplet size if no other information is available.
(ho)max= 0.00128 (tr)o (ΔSG ) d2m / µ Eq.3-5
(ho)max= 0.00128 x 10 (0.194) (500)2 /10
(ho)max= 62.8

3. Calculate Aw/A
Aw/A = 0.5 Qw (tr)w / [ (tr)o Qo + (tr)w Qw ] Eq.3-7
Aw/A = 0.5 x3000 x 10 / [10 x 5000+10 x 3000 ]
Aw/A = 0.1875

4. Determine β from Figure 3-20. With Aw/A = 0.1875, read β = 0.257.

5. Calculate dmax.

dmax = (ho)max/β Eq. 3-8
dmax = 62.8/0.257
dmax = 244 in.

6. Calculate combinations of d, Leff for d less than dmax that satisfy the gas capacity constraint. Use 100-micron droplet size if no other information is available.
dLeff = 420 (TZQ/P) [CD ρc/(ρd- ρc) dm ]0.5 Eq. 3-3
dLeff = 420 ((550 x 0.99 x5)/100) [2.01x 0.3/(54.7- 0.3) 100 ]0.5
dLeff = 120

7. Calculate combinations of d, Leff for d less than dmax that satisfy the oil and water retention time constraints.
d2Leff = 1.42 [(Qw)(tr)w + (Qo)(tr)o] Eq. 3-4
d2Leff = 1.42 [3000 x 10 + 5000 x 10]
d2Leff = 113600
In a table, select different diameters d, and for each d, calculate Leff, and slenderness ratio.

Table 3-3, solution of example 2.

8. Estimate seam-to-seam length.
Lss = Leff + d/12 (gas capacity) Eq. 3-11
Lss = 4/3 Leff (liquid capacity) Eq. 3-12
9. Select slenderness ratio (12Lss/d). Choices in the range of 3 to 5 are common.
10. Choose a reasonable size that does not violate gas capacity constraint or oil-pad thickness constraint. Possible choices are 72 in. diameter by 30 ft. seam-by-seam and 84 in. diameter by 25 ft. seam-by-seam.

3.7: Separator Design (second method)
The following principles of design for liquid-liquid separation apply equally for horizontal or vertical separators.
The settling velocity is a function of gravity and viscosity in accordance with Stoke’s Law. The settling velocity of spheres through a fluid is directly proportional to the difference in densities of the sphere and the fluid, and the square of the diameter of the sphere (droplet), while it is inversely proportional to the viscosity of the fluid.
The liquid-liquid separation capacity of separators may be determined from Equations 3-24 and 3-25.
Values of Cs are found in Table 3-4.
Table 3-1 provides suggested residence times for various liquid-liquid separation applications.
These figures generally assume equal residence times for both the light and heavy liquid phases.
While the residence time approach for liquid-liquid separation equipment design has been widely used in industry for years, it does have some limitations.
• Residence times do not take into account vessel geometry, i.e. 3 minutes residence time in the bottom of a tall, small diameter vertical vessel will not achieve the same separation performance as 3 minutes in a horizontal separator, again according to droplet settling theory.
• The residence time method does not provide any direct indication as to the quality of the separated liquids, e.g. amount of water in the hydrocarbon or the amount of hydrocarbon in the water. Droplet settling theory can not do this either in most cases, but there is some empirical data available which allows for approximate predictions in specific applications.

Vertical Vessel
Wcl = 0.785 Cs D2v (ΔSG)/ µ Eq. 3-24

where
Wcl = flow rate of light condensate liquid, bbl/day
Dv = inside diameter of vessel, ft
ΔSG = difference in sp.gr Kg/l.
µ = viscosity of continuous phase, cp
Cs = Separator Constant, table 3-4

Horizontal Vessel
Wcl = Cs LI HI (ΔSG)/ µ Eq. 3-25

Where
LI = length of liquid interface area, ft.
HI = width of liquid interface area, ft.

Table. 3-4. Values of Cs used in equations 3-24 and 3-25.

The liquid capacity of a separator or the settling volume required can be determined from Eq 3-26 using the retention time given in table 3-1.
V = (W (t))/1440 Eq. 3-26
Where
W = Liquid handling capacity, bbl/day.
V = Liquid settling volume, bbl
t = Retention time, minutes
The following example shows how to size a liquid-liquid separator.

Example 3-3: Determine the size of a vertical separator to handle 600 bpd of 55° API condensate and 50 bpd of produced water. Assume the water particle size is 200 microns. Other operating conditions are as follows:
Operating temperature = 80°F
Operating pressure= 1,000 psig
Water specific gravity = 1.01
Condensate viscosity = 0.55 cp @ 80°F
Condensate specific gravity for 55° API = 0.76

Solution:
From eqn. 3-24
Wcl = 0.785 Cs D2v (ΔSG)/ µ Eq. 3-24

From table 3-4, for free liquids with water particle diameter = 200 microns, Cs = 1,100

600 bbl/day = 0.0785 x 1,100 x [1.01 – 0.76) / 0.55] x D2v = 660 / 392.5 = 1.53 ft2
Dv = 1.24 feet

Using a manufacturer’s standard size vessel might result in specifying a 20-inch OD separator.

Using the alternate method of design based on retention time as shown in Equation 3-26 should give:
V = (W (t))/1440 Eq. 3-26

Vl = ql (t) / 1440
From table 3-1, use 3 minutes retention time
Vl = 650 x 3 / 1440 = 1.35 bbl
Vl = 1.35 x 5.61458 = 7.58 ft3.
Vl = 3.14 * D2 (ft) * h (height of liquid settling room, ft) / 4
Assuming a 24 -inch inlet diameter vessel (2 ft).
7.58 = 3.14 * 4 * h /4
Minimum height of liquid settling room will be h, ft. = 2.4 ft.
Vessel height = +/- (2.4 * 100/30) = 7.5 ft.

It should be remembered that the separator must also be designed for the vapor capacity to be handled. In most cases of high vapor-liquid loadings that are encountered in gas processing equipment design, the vapor capacity required will dictate a much larger vessel than would be required for the liquid load only. The properly designed vessel has to be able to handle both the vapor and liquid loads. Therefore, one or the other will control the size of the vessel used.

Table. 3-5. Standard Horizontal separator sizes.

yasserkassem · Post by **yasserkassem** » Sat Mar 21, 2020 2:52 pm

Chapter 4

Crude Oil Dehydration - Chapter 4 - Part 1

---------------------
Chapter 4 134
Crude oil dehydration 134
4.1: Introduction 134
4.2: Emulsion 134
4.2.1 Energy of Agitation 135
4.2.2 Emulsifying Agents 136
4.2.3: Stability of oil water emulsion 137
4.2.4: Emulsion Treating Theory 139
4.2.5: Demulsifiers 140
4.3: Crude oil treating systems 143
4.3.1: Free-Water Knockouts 143
4.3.2: Gunbarrel tanks with internal and external gas boots 144
4.3.3: Heaters 146
4.4: Emulsion Treating Methods 164
4.4.1: General Considerations 164
4.4.2: Chemical Addition 165
4.5: Heat Required 174
4.5.1: Heat duty 174
4.5.2: Heat Loss 174
4.5.3: Fire Tube Heat Flux 175
4.5.4: Firetube Heat Density 175
4.6: Treater Equipment Sizing 175
4.6.2: Design Procedure 178
4.7: Practical Considerations 184
4.7.1: Gunbarrels with Internal/External Gas Boot 184
4.7.2: Heater-Treaters 184
4.7.3: Electrostatic Heater-Treaters 184

------

Chapter 4

Crude oil dehydration

4.1: Introduction
The fluid produced at the wellhead consists usually of gas, oil, free water, and emulsified water (water–oil emulsion). Before oil treatment begins, we must first remove the gas and free water from the well stream. This is essential in order to reduce the size of the oil–treating equipment.
As presented in previous chapters, the gas and most of the free water in the well stream are removed using separators. Gas, which leaves the separator, is known as ‘‘primary gas.’’ Additional gas will be liberated during the oil treatment processes because of the reduction in pressure and the application of heat. Again, this gas, which is known as ‘‘secondary gas,’’ has to be removed. The free water removed in separators is limited normally to water droplets of 500 µm and larger. Therefore, the oil stream leaving the separator would normally contain free water droplets that are 500 µm and smaller in addition to water emulsified in the oil. This oil has yet to go through various treatment processes (dehydration, desalting, and stabilization) before it can be sent to refineries or shipping facilities.
This chapter deals with the dehydration stage of treatment. The objective of this treatment is first to remove free water and then break the oil emulsions to reduce the remaining emulsified water in the oil. Depending on the original water content of the oil as well as its salinity and the process of dehydration used, oil-field treatment can produce oil with a remnant water content of between 0.2 and 1%. The remnant water is normally called the bottom sediments and water (B.S&W). The treatment process and facilities should be carefully selected and designed to meet the contract requirement for B.S&W. Care should be taken not to exceed the target oil dryness. Removal of more remnant water than allowed by contract costs more money without any benefit.
The basic principles for the treating process are as follows
1. Breaking the emulsion, which could be achieved by either any, or a combination of the addition of heat, the addition of chemicals, and the application of electrostatic field
2. Coalescence of smaller water droplets into larger droplets
3. Settling, by gravity, and removal of free water
4.2: Emulsion
Rarely does oil production takes place without water accompanying the oil. Salt water is thus produced with oil in different forms as follows:
1- Free Water (F.W.)
2- Suspended Water (SS.W)
3- Soluble Water (S.W.)
4- Emulsified Water (E.W.) which may be Oil in water emulsion O/W, or water in oil emulsion W/O.
Apart from free water, emulsified water (water-in-oil emulsion) is the one form that poses all of the concerns in the dehydration of crude oil.
Oil emulsions are mixtures of oil and water. In general, an emulsion can be defined as a mixture of two immiscible liquids, one of which is dispersed as droplets in the other (the continuous phase), and is stabilized by an emulsifying agent. In the oil field, crude oil and water are encountered as the two immiscible phases together. They normally form water-in-oil emulsion (W/O emulsion), in which water is dispersed as fine droplets in the bulk of oil. This is represented in figure 4-1.
However, as the water cut increases, the possibility of forming reverse emulsions (oil-in-water, or O/W emulsion) increases. This is represented in Figure 4-2.
For two liquids to form a stable emulsion, three conditions must exist:
1. The two liquids must be immiscible.
2. There must be sufficient energy of agitation to disperse one phase into the other.
3. There must be the presence of an emulsifying agent.
Conditions 2 and 3 are discussed in the following subsections.

Fig. 4-1 water in oil emulsion.

Fig. 4-2 Oil in water emulsion.
4.2.1 Energy of Agitation
Emulsions normally do not exist in the producing formation, but are formed because of the agitation that occurs throughout the oil production system.
Starting within the producing formation, the oil and water migrate through the porous rock formation, making their way into the wellbore, up the well tubing, through the wellhead choke, and through the manifold into the surface separators. Throughout this journey, the fluids are subjected to agitation due to the turbulent flow. This energy of agitation, which forces the water drops in the bulk of oil, functions in the following pattern:
First, energy is spent to overcome the viscous force between the liquid layers, leading to their separation into thin sheets or parts. This is what we call ‘‘shearing energy’’.
Second, energy is used in the formation of ‘‘surface energy,’’ which occurs as a result of the separation of the molecules at the plane of cleavage. This surface energy is related to the surface tension, which involves the creation of an enormous area of interface with attendant free surface energy. Energy contained per unit area is referred to as ‘‘surface tension,’’.
The drops attain the spherical shape, which involves the least energy contained for a given volume. This is in accordance with the fact that all energetic systems tend to seek the lowest level of free energy. Because the surface tension the volume of a liquid to be contracted or reduced to a shape or a form with the least surface area. This is the reason that causes raindrops to assume a spherical shape.
A schematic presentation of energy utilization in emulsion formation is given in Figure 4-3.

Fig. 4-3 Forms of energy participating in emulsifications
A crucial question that can be asked now is the following:
Can the plant designer prevent emulsion formation? Well, the best he can do is to reduce its extent of formation based on the fact that the liquids initially are not emulsified. From the design point of view, primarily reducing the flowing velocity of the fluid and minimizing the restrictions and sudden changes in flow direction could minimize formation of emulsion.

4.2.2 Emulsifying Agents
If an oil emulsion is viewed through a microscope, many tiny spheres or droplets of water will be seen dispersed through the bulk of oil, as depicted in Figure 4-4.
A tough film surrounds these droplets; this is called a stabilizing film. Emulsifying agents, which are commonly found in crude oil or water in the natural state or introduced in the system as contaminants during drilling and/or maintenance operations, create this film.
These emulsifying agents support the film formation encasing the water droplets, hence the stability of an emulsion.
When studying emulsion stability, it may be helpful to realize that in a pure oil and pure water mixture, without an emulsifying agent, no amount of agitation will create an emulsion. If the pure oil and water are mixed and placed in a container, they quickly separate. The natural state is for the immiscible liquids to establish the least contact or smallest surface area. The water dispersed in the oil forms spherical drops. Smaller drops will coalesce into larger drops, and this will create a smaller interface area for a given volume. If no emulsifier is present, the droplets will eventually settle to the bottom, causing the smallest interface area. This type of mixture is a true “dispersion.”

Fig. 4-4 emulsifier agent film surrounding water droplet
An emulsifying agent in the system is a material, which has a surface active behavior. Some elements in emulsifiers have a preference for the oil, and other elements are more attracted to the water. An emulsifier tends to be insoluble in one of the liquid phases. It thus concentrates at the interface. There are several ways emulsifiers work to cause a dispersion to become an emulsion. The action of the emulsifier can be visualized as one or more of the following:
1. It decreases the interfacial tension of the water droplet, thus causing smaller droplets to form. The smaller droplets take longer to coalesce into larger droplets, which can settle quickly.
2. It forms a viscous coating on the droplets, which keeps them from coalescing into larger droplets when they collide. Since coalescence is prevented, it takes longer for the small droplets, which are caused by agitation in the system, to settle out.
3. The emulsifiers may be polar molecules, which align themselves in such a manner as to cause an electrical charge on the surface of the droplets. Since like electrical charges repel, two droplets must collide with sufficient force to overcome this repulsion before coalescence can occur.

Naturally occurring surface-active materials normally found in crude oil serve as emulsifiers. Paraffins, resins, organic acids, metallic salts, colloidal silts and clay, and asphaltenes (a general term for material with chemical compositions containing sulfur, nitrogen, and oxygen) are common emulsifiers in oil fields. Workover fluids and drilling mud are also sources of emulsifying agents.
The type and amount of emulsifying agent have an immediate effect on the emulsion’s stability. It has been shown that the temperature history of the emulsion is also important as it affects the formation of paraffins and asphaltenes. The speed of migration of the emulsifying agent to the oil–water interface and the behavior in terms of the strength of the interface bond are important factors. An emulsion treated soon after agitation, or soon after the creation of paraffins and asphaltenes, can be less stable and easier to process if the migration of the emulsifier is incomplete. An aged emulsion may become more difficult to treat because the emulsifying agents have migrated to the oil–water interface.
Normally, the lower the crude viscosity and lighter the crude, the more rapid the aging process will be.

4.2.3: Stability of oil water emulsion
The relative difficulty of separating an emulsion into two phases is a measure of its stability. A very stable emulsion is known as a ‘‘tight’’ emulsion and its degree of stability is influenced by many factors.
Accordingly, we can best understand the resolution problem and, hence, the treatment procedure if we consider the following factors:
1- Viscosity of oil: Separation is easier for a less viscous oil phase.
2- Density or gravity difference between oil and water phases: Better separation is obtained for a larger difference.
3- Interfacial tension between the two phases (which is related to the type of emulsifying agent): Separation is promoted if this force is lowered (i.e., decreasing the interfacial tension).
4- Size of dispersed water droplets: The larger the size of water drops, the faster is the separation. This could be readily concluded from Stokes’ law presented in previous chapters, where the velocity of settling is directly proportional to the difference in specific gravity, the square power of droplet diameter, and inversely proportional to the viscosity.
Vt = 1.78 x 10-6 (ΔSG) d2m / µ Eq. 2-10
The size of dispersed water droplets is an important factor in emulsion stability. A typical droplet size distribution for emulsion samples was determined by using a special computer scanning program. Results reported in Figure 4-5 indicate that most of the droplets found in oil emulsions are below 50 um.

Fig. 4-5 Water droplet size distribution in water in oil emulsion.

5- Percentage of dispersed water: The presence of a small percentage of water in oil under turbulence conditions could lead to highly emulsified mixture. Water droplets are finely divided and scattered with very little chance of agglomerating to larger particles.
6- Salinity of emulsified water: Highly saline water will lead to a faster separation because of a higher density difference between the oil and the water phases.
7- Age of the Emulsion: As emulsions age they become more stable and separation of the water droplets becomes more difficult. The time required to increase stability varies widely and depends on many factors. Before an emulsion is produced, the emulsifying agents are evenly dispersed in the oil. As soon as the water phase is mixed with the oil, the emulsifying agents begin to cluster around the water droplet to form a stable emulsion. While the initial stabilization may occur in a matter of a few seconds, the process of film development may continue for several hours. It will continue until the film around the droplet of water is so dense that no additional stabilizer can be attracted, or until no stabilizer is left to be extracted from the oil. At such a time the emulsion has reached a state of equilibrium and is said to be aged. The older the emulsion, the more difficult it is to treat. Therefore, emulsion breaking or treating operations are often located as close to the wellhead as possible, so that emulsions formed during flow in the production tubing and wellhead equipment are not allowed to age before treatment.
8- Presence and Concentration of Emulsifying Agents : Chemicals (demulsifiers) are normally used to reduce the interfacial tension. Chemical effectiveness is enhanced by mixing, time, and temperature.
Adequate mixing and sufficient time are required to obtain intimate contact of the chemical with the dispersed phase. A certain minimum temperature is required to ensure the chemical accomplishes its function.
Both viscosity reduction and effectiveness of chemical are dependent on the attainment of a certain minimum temperature. It may well be that the increase in chemical effectiveness is a result of the decrease in viscosity of the oil phase.
9- Agitation: The type and severity of agitation applied to an oil–water mixture determine the water drop size. The more turbulence and shearing action present in a production system, the smaller the water droplets and the more stable the emulsion will be.
Emulsions are formed during production of the fluid. The degree of emulsification is dependent on the agitation of the two phases by pumps, chokes, etc.

The above factors determine the “stability” of emulsions. Some stable emulsions may take weeks or months to separate if left alone in a tank with no treating. Other unstable emulsions may separate into relatively clean oil and water phases in just a matter of minutes.

4.2.4: Emulsion Treating Theory
Removing water from crude oil often requires additional processing beyond the normal oil–water separation process, which relies on gravity separation. Crude oil treating equipment is designed to break emulsions by coalescing the water droplets and then using gravity separation to separate the oil and water. In addition, the water droplets must have sufficient time to contact each other and coalesce. The negative buoyant forces acting on the coalesced droplets must be sufficient to enable these droplets to settle to the bottom of the treating vessel. Therefore, it’s important when designing a crude oil treating system to take into account temperature, time, viscosity of the oil, which may inhibit settling, and the physical dimensions of the treating vessel, which determines the velocity at which settling must occur.
When selecting a treating system, several factors should be considered to determine the most desirable method of treating the crude oil to contract requirements. Some of these factors are
• Stability (tightness) of the emulsion,
• Viscosity of crude oil and emulsion at different temperature,
• Specific gravity of the oil and produced water,
• Corrosiveness of the crude oil, produced water, and associated gas,
• Scaling tendencies of the produced water,
• Quantity of fluid to be treated and percent water in the fluid,
• Paraffin-forming tendencies of the crude oil, and pour point of crude oil.
• Desirable operating pressures for equipment,
• Availability of a sales outlet and value of the associated gas produced.
A common method for separating this “water-in-oil” emulsion is to heat the stream. Increasing the temperature of the two immiscible liquids deactivates the emulsifying agent, allowing the dispersed water droplets to collide. As the droplets collide they grow in size and begin to settle.
If designed properly, the water will settle to the bottom of the treating vessel due to differences in specific gravity.
Laboratory analysis, in conjunction with field experience, should be the basis for specifying the configuration of treating vessels. The purpose of this chapter is to present a rational alternative for those instances when laboratory data do not exist or, if it is desirable, to extrapolate field experience.

4.2.5: Demulsifiers
4.2.5.1: Introduction
Emulsions can be resolved or broken thermally and/or chemically. When we chemically resolve an emulsion, we use a demulsifier or emulsion breaker. These two names are used interchangeably and describe the same chemical. Chemical demulsifiers sold under various trade names, such as Tretolite, Visco, Breaxit, etc., are highly useful in resolving emulsions.
Demulsifiers act to neutralize the effect of emulsifying agents. Typically, they are surface-active agents and thus their excessive use can decrease the surface tension of water droplets and actually create more stable emulsions. In addition, demulsifiers for water-in-oil emulsions tend to promote oil-in-water emulsions; therefore, excessive chemical use may cause water treating problems.
Four important actions are required of a demulsifier:
• Strong attraction to the oil–water interface,
• Flocculation,
• Coalescence,
• Solid wetting.
When these actions are present, they promote the separation of oil and water. The demulsifier must have the ability to migrate rapidly through the oil phase to the droplet interface, where it must compete with the more concentrated emulsifying agent. The demulsifier must produce an attraction for similar droplets. In this way large clusters of droplets gather, which, under a microscope, appear like bunches of fish eggs.
If the emulsifier is weak, the flocculation force may be enough to cause coalescence. This is not true in most cases, and the demulsifier must therefore neutralize the emulsifier and promote a rupture of the droplet interface film. This is the opener that causes coalescence. With the emulsion in a flocculated condition, the film rupture results in rapid growth of water drop size.
The manner in which the demulsifier neutralizes the emulsifier depends upon the type of emulsifiers. Iron sulfides, clays, and drilling muds can be water wet, causing them to leave the interface and be diffused into the water droplet. Paraffins and asphaltenes could be dissolved or altered to make their films less viscous so they will flow out of the way on collision or could be made oil wet so they will be dispersed in the oil.
It would be unusual if one chemical structure could produce all four desirable actions. A blend of compounds is therefore used to achieve the right balance of activity.
The demulsifier selection should be made with the process system in mind. If the treating process is a settling tank, a relatively slow-acting compound can be applied with good results. On the other hand, if the system is a chemical-electric process where some of the flocculation and coalescing action is accomplished by an electric field, there is need for a quick-acting compound.
As field conditions change, the chemical requirements can change.
If the process is modified, e.g., very low rates on electrostatic units, the chemical requirements can change. Seasonal changes bring paraffin induced emulsion problems. Workovers contribute to solids and acid/base contents, which alters the emulsion stability. So no matter how satisfactory a demulsfier is at one point in time, it may not be satisfactory over the life of the field.
The cost to dehydrate crude oil chemically is a function of several factors.
First, the ratio of oil to water is important—it is generally easier and, hence, less costly to dehydrate crudes with very high water cuts. Next, the severity of the emulsion is important. A “tight” emulsion consisting of small droplets is much more difficult to break—it has a higher surface area to volume ratio than a “loose” emulsion and, hence, the demulsifier has more work to do to seek out the interface. Next, the residence time available for separation is important. Small residence times inhibit complete separation of water droplets from oil. This may lead to re-entrainment of water as the crude goes from one processing stage to another. The result is ineffective dehydration. Next, Higher temperatures result in lower oil phase viscosities, which enable the demulsifier to migrate to the oil–water interface faster and for coalesced water droplets to drop out easier. Last, the dehydration cost is directly influenced by chemical selection. Poor chemical selection will result in a non-optimized treatment, which will mean higher costs. Chemical selection is not a simple process—it is best left to suppliers. However, one can assist in the process by providing on-site testing opportunities for chemical suppliers to select the best chemicals for specific applications.

Fig. 4-6. Demulsifier action on water droplets.

Fig. 4-7. Droplet growth steps.

4.2.5.2: Bottle Test
This is one of the most common of all the chemical selection tests. Emulsion-breaking chemicals are most commonly tested with a bottle test, which involves mixing various chemicals with samples of the emulsion and observing the results. Such tests are effective in eliminating some chemicals and selecting those that appear to be more efficient. Bottle tests also provide an estimate of the amount of chemical required and an estimate of the settling time required for a treating vessel.
Bottle tests should be performed on a representative sample as soon as the sample is obtained because of the possible detrimental effects of aging.
These tests should be performed at conditions that are as close to field treating conditions as possible. Synthetic water should not be used in place of produced water in bottle tests because the produced water may have very different properties, and it may contain impurities that are not present in the synthetic water.
While candidate chemicals and approximate dosages can be determined in bottle tests, the dynamic nature of the actual flowing system requires that several candidates be field-tested. In actual conditions, the emulsion undergoes shearing through control valves, coalescence in flow-through pipes, and changes to the emulsion that occur inside the treating vessel as a result of inlet diverters, water-wash sections, etc. Static bottle tests cannot model these dynamic conditions.
As well as determining the potential dehydration performance of a demulsifier, the bottle test can also be used to investigate chemical incompatibilities. Here, the performance of a demulsifier is evaluated on a chemical-free sample and then on a sample of crude, which includes the other production chemicals at their respective dose rates. The change in performance, if any, is recorded and the chemical discarded if incompatibilities exist. Another aspect of incompatibility may also be determined, namely, in which order the chemicals should be injected. If the bottle tester is experienced, this order of injection, which will produce subtle changes in the bottle test results, can be investigated and an optimum injection order determined.

4.2.5.3: Field Trial
Having selected a promising demulsifier candidate, a field trial should be carried out to test the chemical’s ability to operate in a dynamic system.
In the field test, the flexibility of the demulsifier to process changes can be established. This data will be useful when the chemical is used in full scale operation. In most field trial situations, the demulsifier being tested is first used in conjunction with a test separator system. This enables the supplier to look at the response of the chemical to one or more wells and to provide the tester an idea of the true field dosage. If this preliminary scenario is successful, the chemical can then be dosed into the full system and optimized for different well configuration and flow rates. In the field trial, the chemical’s response to system upsets can be determined and, hence, an operating response can be set.

4.2.5.4: Field Optimization
After a successful field trial, a full-scale field optimization is carried out.
Here, the chemical performance is monitored routinely as are the possible side effects of under-or overdosing, such as separator interface buildup. It may be that if the field produces through two or more platforms, injection locations and dose rates may need to be optimized for each location.

4.2.5.5: Changing the Demulsifier
As crude characteristics change over the life of a field, the performance of the demulsifier chemical will change also. Typically, when fields first produce water, the emulsions formed are difficult to break. As the field ages and the water cut increases, the stability of the emulsion and even the emulsifying agents themselves may change. Hence, it is usual to investigate demulsifier performance every 2 to 3 years. In some cases where a step change in water cut is experienced, it may be prudent to investigate demulsifier performance more frequently. In most cases a quick bottle test is all that is required to determine if the current chemical is still optimum. If not, a full bottle test to find a more effective or cheaper chemical can be undertaken.

4.2.5.6: Demulsifier Troubleshooting
The most common problem with demulsifiers is overdosing. Poor treatment, dirty water, and interface pad build-up are all symptoms of overdosing an optimum chemical. Overdosing can occur by a step increase in dose rate, e.g., going from 5 to 20 ppm, or by a gradual accumulation of chemical in the system. The latter is most often seen in high water cut systems where a small change from optimum can result in dirty water. The gradual accumulation of chemical usually occurs at the separator interface and is often difficult to detect. However, highly variable water quality caused by intermittent interface sloughing is often a clue to this scenario.
Other problems with demulsifiers can be that their viscosity changes with temperature. Most demulsifiers are viscous chemicals whose ability to be pumped can drop dramatically with reduced temperature. If this is the case, it may be prudent to ask the chemical supplier to produce a “winterized” version of the chemical. This is often done by reducing the percentage of active ingredient and adding more solvent carrier. If this is the solution, the dose rates will need to be re-optimized for best performance.
Another common problem with demulsifiers is their apparent lack of treatment “range.” It is not uncommon for a field demulsifier to have a different performance standard for different wells within a field. In some cases “rogue wells” may exist, which are basically untreatable by the optimum demulsifier for the rest of the system. In these cases two demulsifiers may be used or the original demulsifier may be injected at a higher dose rate or even downhole in the rogue well. The bottle test will often indicate rogue wells and their best treatment solution.
Incompatibility of demulsifiers and corrosion inhibitors are often the cause of poor dehydration performance. Corrosion inhibitors are surfactant chemicals that often act as emulsifying agents, thus making the demulsifier work harder.
In cases of conflict, it is usually easier to blend a new demulsifier or change the injection points of the chemicals. However, in some fields the opposite was true. Corrosion inhibitor replacement was the best way to deal with the incompatibility problem.
As there are no online analyzers for demulsifier performance, one must monitor the facilities for changes in water or crude quality that may be attributed to poor demulsifier performance. Chemical suppliers can help here by giving us the anticipated system response to incompatibilities and over- or under-dosing. They should get this information from the bottle test and the demulsifier field trial.
4.3: Crude oil treating systems
4.3.1: Free-Water Knockouts
Most well streams contain water droplets of varying size. If they collect together and settle to the bottom of a sample within 3 to 10 minutes, they are called free water. This is an arbitrary definition, but it is generally used in designing equipment to remove water that will settle out rapidly. Figure 4-8.
A free-water knockout (FWKO) is a pressure vessel used to remove free water from crude oil streams.
Free water knock out drums are usually located in the production flow path where turbulence has been minimized. Restrictions such as orifices, chokes, throttling globe valves, and fittings create turbulence in the liquids that aggravate emulsions. Free water, at wellhead conditions, frequently will settle out readily to the bottom of an expansion chamber.

Fig. 4-8. Free water knock out (FWKO)

Sizing for these vessels were discussed in previous chapters. Factors affecting design include retention time, flow rate (throughput), temperature, oil gravity (as it influences viscosity), water drop size distribution, and emulsion characteristics. Abnormal volumes of gas in the inlet stream may require proportionately larger vessels as these gas volumes affect the throughput rate. A simple “field check” to determine retention time is to observe a fresh sample of the wellhead crude and the time required for free water to segregate.
In installations where gas volumes vary, a two-phase separator is usually installed upstream of the free-water knockout. The two-phase separator removes most of the gas and reduces turbulence in the free-water knockout vessel. The free-water knockout usually operates at 50 psig or less. Internals should be coated or protected from corrosion since they will be in constant contact with salt water.

4.3.2: Gunbarrel tanks with internal and external gas boots
The gunbarrel tank, sometimes called a wash tank, is the oldest equipment used for multiwell onshore oil treating in a conventional gathering station or tank battery. Gunbarrel tanks are very common in heavy crude applications, and for low flow rate onshore applications for all crude gravities.
The gunbarrel tank is a vertical flow treater in an atmospheric tank.
Figure 4-9A shows a “gunbarrel” tank with an internal gas boot. Typically, gunbarrels have an internal gas separating chamber or “gas boot” extending 6 to 12 ft above the top of the tank, where gas is separated and vented, and a down-comer extending 2 to 5 ft from the bottom of the tank.

Figure 4-9A illustrates a simple gunbarrel “washing tank” configuration. Fig 4-9B illustrates a gunbarrel configuration with an “external” gas boot. This configuration is preferred on larger tanks, generally in the 60,000-barrel range, where attaching an internal gas boot is structurally difficult. In either case, the gunbarrel tank is nothing more than a large atmospheric settling tank that is higher than downstream oil shipping and water clarifier tanks. The elevation difference allows gravity flow into the downstream vessels.
The emulsion, flowing from an upstream separator and possibly a heater, enters the top of the gas separation section of the gas boot. The gravity separation section removes flash gas and gas liberated as a result of heating the emulsion. The emulsion flows down the down-comer to a spreader, which is positioned below the oil–water interface. The emulsion rises to the top of the surrounding layer of water. The water level is controlled by a water leg or automatic level control. The emulsion passage through the water helps collect the entrained water and converts the emulsion into distinct oil and water layers. Oil accumulates at the top and flows out through the spillover line into the oil-settling tank. Water flows from the bottom of the tank, up through the water leg, and into a surge or clarifier tank. The settling time in the vessel for the total fluid stream is usually 12 to 24 hr.
Most gunbarrels are unheated, though it is possible to provide heat by heating the incoming stream external to the tank, installing heating coils in the tank, or circulating the water to an external or “jug” heater in a closed loop. It is preferable to heat the inlet so that more gas is liberated in the boot, although this means that fuel will be used in heating any free water in the inlet.
The difference in height between the oil spillover line and the external water leg controls the oil-water interface.

Example 4-1: Determination of external water leg height
Given:
Oil gravity @ 600F = 360API,
Water specific gravity = 1.05,
Height of oil outlet = 23 ft,
Height of interface level = 10 ft (for this example),
Height of water outlet = 1 ft,

Solution:
(1) Determine the oil specific gravity.
Oil specific gravity = 141.5/ (131.5+ 0API)
= 141.5/ (131.5+36) = 0.845
(2) Determine the oil gradient.
Since the change in the pressure with depth for fresh water is 0.433 psi/ft of depth, the change in pressure with depth of fluid whose specific gravity is SG would be [0.433 x (SG)]; thus, the oil gradient is Oil gradient = 0.433 x 0.845 = 0.366 psi/ft.
(3) Determine the water gradient.
Water gradient = 0.433 * 1.05 = 0.455 psi/ft.
(4) Calculate the height of the oil and the height of the water in the tank.
Ho = Height of oil outlet − height of interface level = 23−10 = 13 ft.
Hw = Height of interface level − height of water outlet = 10−1 = 9 ft.
(5) Perform a pressure balance.
Hydrostatic Pressure inside Tank = Hydrostatic Pressure in the Water Leg
(13 x 0.366) + (9 x 0.455) = H x 0.455
H = 19.5 ft.

The design details for the spreader, water leg, and gas separation section vary for different manufacturers. These details do not significantly affect the sizing of the tank.
No matter how careful the design of the spreaders, large wash tanks are very susceptible to short-circuiting. This is due to temperature and density differences between the inlet emulsion and the fluid in the tank, solids deposition, and corrosion of the spreaders.
Short-circuiting means that: The inlet fluid does not mix well with existing fluid due to difference in temperature, and it moves to the outlet without spending the normal residence time.

Standard tank dimensions are listed in API Specification 12F (Shop Welded Tanks), API Specification 12D (Field Welded Tanks), and API Specification 12B (Bolted Tanks).

Gunbarrels are simple to operate and, despite their large size, are relatively inexpensive. However, they have a large footprint, which is why they are not used on offshore platforms. Gunbarrels hold a large volume of fluids, which is a disadvantage should a problem develop.

Figure 4-9A. Gunbarrel with blanketing gas for low gas oil ratio crude.

When the treating problem is detected in the oil outlet, a large volume of bad oil has already collected in the tank. This oil may have to be treated again, which may require large slop tanks, recycle pumps, etc. It may be beneficial to reprocess this bad oil in a separate treating facility so as to avoid further contamination of the primary treating facility.

Gunbarrels are most often used in old, low-flow-rate, onshore facilities.
In recent times, vertical heater-treaters have become so inexpensive that they have replaced gunbarrels in single-well installations. On larger installations onshore in warm weather areas, gunbarrels are still commonly used. In areas that have a winter season it tends to become too expensive to keep the large volume of oil at a high enough temperature to combat potential pour-point problems.

4.3.3: Heaters
Heaters are vessels used to raise the temperature of the liquid before it enters a gunbarrel, or any further treating vessels used to treat crude oil emulsions. The two types of heaters commonly used in upstream operations are indirect fired heaters and direct fired heaters.
Both types have a shell and a fire tube. The fire tube contains within it a flame caused by the mixture of air and natural gas ignited by a pilot light, and the hot exhaust gases which result from this combustion. The hot external surface of the fire tube heats a bath of liquid in which it is immersed.
Indirect heaters have a third element, which is the process flow coil. Heaters have standard accessories such as burners, regulators, relief valves, thermometers, temperature controllers, etc.

Figure 4-9B. Gunbarrel with an internal gas boot.

4.3.3.1: Indirect Fired Heaters
Figures 4-10, 4-11, and 4-12, show indirect fired heater (bath heater). Oil flows through tubes that are immersed in water, which in turn is heated by a fire tube.
Alternatively, heat may be supplied to the water bath by a heating fluid medium, steam, or electric immersed heaters instead of a fire tube. Indirect heaters maintain a constant temperature over a long period of time and are safer than direct heaters.
Hot spots are not as likely to occur on the fire tube if the calcium content of the heating water is controlled. The primary disadvantage is that these heaters require several hours to reach the desired temperature after they have been out of service.

4.3.3.2: Direct Fired Heaters
Figure 4-13 shows a typical direct fired heater. Oil flows through an inlet distributor and is heated directly by a fire box. Alternatively, heat may be supplied to by a heating fluid medium, steam, or an electric immersed heater instead of the fire tube. Direct fired heaters are quick to reach the desired temperature, are efficient (75 to 90%), and offer a reasonable initial cost. Direct fired heaters are typically used where fuel gas is available and high volume oil treating is required. On the other hand, they are hazardous and require special safety equipment. Scale may form on the oil side of the fire tube, which prevents the transfer of heat from the fire box to the oil emulsion. Heat collects in the steel walls under the scale, which causes the metal to soften and buckle. The metal eventually ruptures and allows oil to flow into the fire box, which results in a fire. The resultant blaze, if not extinguished, will be fed by the incoming oil stream.

4.3.3.3: Waste Heat Recovery
A waste heat recovery heater captures waste heat from the exhaust stacks of compressors, turbines, generators, and large engines. These hot exhaust gases can be routed through a tube and immersed in a bath performing the same function as a fire tube. Alternatively, heat exchangers may be used to transfer this heat to a heating fluid medium, which in turn is used to heat the crude oil emulsion.

Fig.4-10. Indirect fired heater.

Fig.4-11. Indirect fired heater.

Fig. 4-12. API 12K Indirect fired heater assembly.

Fig. 4-13. Direct fired heater assembly.
4.3.3.4: Heater-Treaters
Heater-treaters are an improvement over the gunbarrel and heater system.
Many designs are offered to handle various conditions such as viscosity, oil gravity, high and low flow rates, corrosion, and cold weather. When compared to gunbarrels, heater-treaters are less expensive initially, offer lower installation costs, provide greater heat efficiency, provide greater flexibility, and experience greater overall efficiency. On the other hand, they are more complicated, provide less storage space for basic sediment, and are more sensitive to chemicals. Since heater-treaters are smaller than other treating vessels, their retention times are minimal (10 to 30 min).
Build-up of sediment on the walls or bottom of the treater can cause reduction on retention time, and cause the interface levels to rise and liquid to carry over and/or oil to exit the treater with salt water. An annual inspections should be performed to include internal inspection for corrosion, sediment build-up, and scale build-up.

4.3.3.4.1: Vertical Heater-Treaters
The most commonly used single-well treater is the vertical heater-treater, which is shown in Figure 4-14. The vertical heater-treater consists of four major sections: gas separation, free-water knockout, heating and water wash, and coalescing-settling sections. Incoming fluid enters the top of the treater into a gas separation section, where gas separates from the liquid and leaves through the gas line. Care must be exercised to size this section so that it has adequate dimensions to separate the gas from the inlet flow. If the treater is located downstream of a separator, the gas separation section can be very small. The gas separation section should have an inlet diverter and a mist extractor.
The liquids flow through a down-comer to the base of the treater, which serves as a free-water knockout section. If the treater is located downstream of a free-water knockout or a three-phase separator, the bottom section can be very small. If the total well stream is to be treated, this section should be sized for 3 to 5 minutes’ retention time to allow the free water to settle out.
This will minimize the amount of fuel gas needed to heat the liquid stream rising through the heating section. The end of the down-comer should be slightly below the oil–water interface so as to “water-wash” the oil being treated. This will assist in the coalescence of water droplets in the oil.
The oil and emulsion rise through the heating and water-wash section, where the fluid is heated (Figure 4-15). A fire tube is commonly used to heat the emulsion in the heating and water-wash section.

After the oil and emulsion are heated, the heated oil and emulsion enter the coalescing section, where sufficient retention time is provided to allow the small water droplets in the oil continuous phase to coalesce and settle to the bottom. As shown in Figure 4-16, baffles are sometimes installed in the coalescing section to treat difficult emulsions.
The baffles cause the oil and emulsion to follow a back-and-forth path up through the treater. Treated oil flows out the oil outlet, at the top of the coalescing section, and through the oil leg heat exchanger, where a valve controls the flow. Heated clean oil preheats incoming cooler emulsion in the oil leg heat exchanger (Figure 4-17). Separated water flows out through the water leg, where a control valve controls the flow to the water treating system (Figure 4-16).
As shown in Figure 4-18, any gas, flashed from the oil due to heating, is captured on the condensing head. Any gas that didn’t condense flows through an equalizing line to the gas separation section.
As shown in Figure 4-19, a vane-type mist extractor removes the liquid mist before the gas leaves the treater. The gas liberated when crude oil is heated may create a problem in the treater if it is not adequately designed. In vertical heater-treaters the gas rises through the coalescing section. If a great deal of gas is liberated, it can create enough turbulence and disturbance to inhibit coalescence. Equally important is the fact that small gas bubbles have an attraction for surface-active material and hence water droplets. Thus, they tend to keep the water droplets from settling out and may even cause them to carry over to the oil outlet.
The oil level is maintained by pneumatic or lever-operated dump valves. The oil–water interface is controlled by an interface level controller or an adjustable external water leg.

Standard vertical heater-treaters are available in 20- and 27-ft heights. These heights provide sufficient static liquid head so as to prevent vaporization of the oil.

Figure 4-14. Schematic of a vertical heater-treater.

Figure 4-15. 3d view illustrating oil and emulsion rising through the heating and water-wash.

Figure 4-16. Baffles, installed in the coalescing section, cause the emulsion to follow a back-and-forth path up through the oil settling section (left). Vertical heater treater with a water leg (right).
.

Figure 4-17. Heated clean oil preheats incoming cooler emulsion in the oil leg heat exchanger

Figure 4-18. Gas, flashed from the oil during heating, is captured on the condensing head.

Figure 4-19. Vane-type mist extractor removes the liquid mist before the gas leaves the treater
.

Fig.4-20. API 12L. Vertical heater treater assembly.
Coalescing Media
It is possible to use coalescing media to promote coalescence of the water droplets. These media provide large surface areas upon which water droplets can collect. In the past the most commonly used coalescing media was wood shavings or “excelsior,” which is also referred to as a “hay section.” The wood excelsior was tightly packed to create an obstruction to the flow of the small water droplets and promote random collision of these droplets for coalescence. When the droplets were large enough, they fell out of the flow stream by gravity. Figure 4-22 shows a vertical heater-treater utilizing an excelsior section.
The use of an excelsior section allowed lower treating temperatures. However, these media had a tendency to clog with time and were difficult to remove. Therefore, they are no longer used.

Figure 4-22. Excelsior in vertical heater treater aids in coalescence of water droplets.

4.3.3.4.2: Horizontal Heater-Treaters
For most multiwell flow streams, horizontal heater-treaters are normally required. Figure 4-23 shows a simplified schematic of a typical horizontal heater-treater. Design details vary from manufacturer to manufacturer, but the principles are the same. The horizontal heater-treater consists of three major sections: front (heating and water-wash), oil surge chamber, and coalescing sections.(fig. 4-24 shows API 12L. Horizontal heater treater assembly).

As shown in Figures 4-23 and 4-25, the oil, emulsion, and free water pass around the deflector hood to the spreader located slightly below the oil–water interface, where the liquid is “water-washed” and the free water is separated. For low gas–oil-ratio crudes, blanket gas may be required to maintain gas pressure. The oil and emulsion are heated as they rise past the fire tubes and are skimmed into the oil surge chamber.
As free water separates from the incoming fluids in the front section, the water level rises. If the water is not removed, it will continue to rise until it displaces all emulsion and begins to spill over the weir into the surge section (In the same time, heat loss will increase due to heating the water instead of heating the emulsion for treatment). On the other hand, if the water level becomes too low, the front section will not be able to water-wash the incoming oil and emulsion, which would reduce the efficiency of the treater. Therefore, it is important to accurately control the oil–water interface in the front section.

Fig. 4-23. Simplified horizontal heater treater.

Fig.4-24. API 12L. Typical horizontal heater treater assembly.

Figure 4-25. Horizontal heater-treater flow pattern.

Incoming fluids enter the front (heating and water-wash) section through the fluid inlet and down over the deflector hood (Figures 4-23 and 4-25) where gas is flashed and removed. Heavier materials (water and solids) flow to the bottom while lighter materials (gas and oil) flow to the top. Free gas breaks out and passes through the gas equalizer loop to the gas outlet.

Figure 4-26. Horizontal heater-treater showing the oil, emulsion, and free water passing around the deflector hood to the spreader located slightly below the oil–water interface where the liquid is “water-washed” and the free water separated.
The oil–water interface is controlled by an interface level controller, which operates a dump valve for the free water (refer to Figure 4-27).
As shown in Figure 4-28, a level safety low shutdown sensor is required in the upper portion of the front (heating and water-wash) section. This sensor assures liquid is always above the fire tube. If the water dump valve malfunctions or fails open, the liquid surrounding the fire tube will drop, thus not absorbing the heat generated from the fire tube and possibly damaging the fire tube by overheating. Thus, if the level above the fire tube drops, the level safety low shutdown sensor sends a signal that closes the fuel valve feeding the fire tube. It is also important to control the temperature of the fluid in the front (heating and water-wash) section.
Therefore, a temperature controller, controlling the fuel to the burner or heat source, is required in the upper part of the heating–water-wash section (refer to Figure 4-29).

Figure 4-27. Oil–water interface in the heating and water-wash section is controlled by an interface level controller.

Figure 4-28. Level safety low sensor, located at the top of the heating–water-wash section, shuts off the fuel to the heat source (fire-tube) on low liquid level.

Figure 4-29. Temperature controller, located in the upper part of the heating–water-wash section, controls the fuel to the burner or heat source.

Figure 4-30. Level controller in the oil surge section operates the clean oil dump valve.

A level controller, in the oil surge section (refer to Figure 4-30), operates the dump valve on the clean oil outlet line. This dump valve regulates the flow of oil out the top of the vessel, which maintains a liquid packed condition. When the clean oil outlet valve is open, the pressure of the gas in the surge chamber forces the emulsion to flow through the spreader and push the clean oil through the clean oil collector (Figure 4-31). When the clean oil outlet valve closes, the flow of emulsion to the coalescing-settling section stops since the coalescing-settling section is completely full of liquid.
The oil and emulsion flow through a spreader into the back or coalescing section of the vessel, which is fluid packed. The spreader distributes the flow evenly throughout the length of this section. Because it is lighter than the emulsion and water, treated oil rises to the clean oil collector, where it is collected and flows to the clean oil outlet. The collector is sized to maintain uniform vertical flow of the oil. Coalescing water droplets fall countercurrent to the rising oil continuous phase.
The front (heating and water-wash) section must be sized to handle settling of the free water and heating of the oil. The coalescing section must be sized to provide adequate retention time for coalescing to occur and to allow the coalescing water droplets to settle downward countercurrent to the upward flow of the oil.

Most horizontal heater-treaters built today do not use fire tubes. Heat is added to the emulsion in a heat exchanger before the emulsion enters the treater. In these cases the inlet section of the treater can be fairly short because its main purpose is to degas the emulsion before it flows to the coalescing section.
Some heater-treaters are designed with only the coalescing section. In these cases the inlet is pumped through a heat exchanger to a treater that operates at a high enough pressure to keep the oil above its bubble point. Thus, the gas will not evolve in the coalescing section of the treater.

Figure 4-31. Pressure of the gas in the surge section forces the emulsion to flow through the spreader in the coalescing section and push the clean oil out through the clean oil collector.

yasserkassem · Post by **yasserkassem** » Sun Mar 22, 2020 1:29 pm

Chapter 4 - Part 2

4.3.3.4.3: Electrostatic Heater-Treaters
Some horizontal heater-treaters add an electrostatic grid in the coalescing section. Figure 4-32A and 4-32B illustrate typical horizontal electrostatic treaters.
The flow path in an electrostatic heatertreater is basically the same as in a horizontal heater-treater, except that an electrostatic grid is included in the coalescing-settling section, which helps to promote coalescence of the water droplets.
The electrostatic section contains two or more electrodes. An electrical system supplies an electric potential to the electrodes. The usual applied voltage ranges from 10,000 to 35,000 VAC, and the power consumption is from 0.05 to 0.10KVA/ft2 (0.54 to 1.08KVA/m2) of grid. The intensity of the electrostatic field is controlled by the applied voltage and spacing of electrodes. In some installations the location of the ground electrode can be adjusted externally to increase or decrease its spacing to the “hot” electrode. The use of an electric field is most effective whenever the fluid viscosity is less than 50 cp at separating temperature, the specific gravity difference between the oil and water is greater than 0.001, and the electrical conductivity of the oil phase does not exceed 10-6 mho/cm.
The electrical control system that supplies energy to the electrodes consists of a system of step-up transformers (either single or three phase) in which the primary side is connected to a low-voltage power source (220 to 440 V) and secondary windings are designed so that the induced voltage will be of the desired magnitude (refer to Figure 4-33).

Figure 4-32A. Horizontal electrostatic heater-treater.

Figure 4-32B. Horizontal electrostatic heater-treater.

As shown in Figure 4-34, oil and small water droplets enter the coalescing section and travel up into the electrostatic grid section, where the water droplets become “electrified” or “ionized” and are forced to collide.
The electrodes have electrical charges that reverse many times a second; thus, the water droplets are placed in a rapid back-and-forth motion. The greater the motion of the droplets, the more likely the water droplets are to collide with each other, rupture the skin of the emulsifying agent, coalesce, and settle out of the emulsion. Because of the forced collisions, electrostatic heater-treaters typically operate at lower temperatures and use less fuel than horizontal heater-treaters. The time in the electronic field is controlled by electrode spacing and the vessel configuration. An electronic field exists throughout the body of the oil within the vessel, even though most coalescing takes place in the more intense fields in the vicinity of the electrodes.

Figure 4-33. Electrical control system of an electrostatic heater-treater.

It is imperative that the design of the vessel provide for distribution of the emulsion across the electrical grid. It is also essential to maintain the fluid in the liquid phase in the electrical coalescing section. Gas evolving in the coalescing section will attract the small water droplets in the emulsion, becoming saturated with water and carrying the water up to the oil outlet. In addition, water-saturated vapors, which are highly conductive, will greatly increase the electrical power consumption.
It is also important to prevent the water level from reaching the height of the electrodes. Nearly all produced water contains some salt. These salts make the water a very good conductor of electric currents. Thus, if the water contacts the electrodes, it may short out the electrode grid or the transformer.
Since coalescence of the water droplets in an electric field is dependent on the characteristics of the specific emulsion being treated, sizing of grid area requires laboratory testing. Field experience tends to indicate that electrostatic treaters are efficient at reducing water content in the crude to the 0.1 to 0.5 percent level. This makes these treaters particularly attractive for desalting operations.

Figure 4-34. Effect of electrical charge on small water droplets in the emulsion.

Oil Dehydrators (Desalter)
The primary factor when designing coalescing units is the loading rate.
Vessels are sized for a certain volume flow per unit time per square foot of grid area. Since coalescence of water droplets in an electric field is so dependent on the characteristics of the particular emulsion to be treated, it is unlikely that a general relationship of water droplet size to use in the settling equations can be developed. Electrostatic treaters are particularly attractive for oil desalting for normal crude treating, where 0.5 to 1.0% BS&W is acceptable, it is recommended that the vessel be sized as a horizontal heater-treater, neglecting any contribution from the electrostatic grids. By trial and error after installation, the electric grids may be able to allow treating to occur at lower temperatures or higher flow rates.

Figure 4-35 shows one variation of the electrostatic heater-treater where the vessel only contains the coalescing section with the electrostatic grid. Units configured in this manner are called “oil dehydrators.” These vessels must have separate upstream vessels for de-gassing, free-water removal and heating. This configuration should be considered when the volume to be treated exceeds 15,000 to 20,000 barrels per day.
4.4: Emulsion Treating Methods
4.4.1: General Considerations
Treating processes and equipment should not be selected until the physical characteristics of the oil and water have been determined and a study of the effect of available chemicals on the emulsion has been made. The water remaining in the crude after the free water has settled out is considered to be in an emulsified state. Emulsified oil is removed by one or more treating processes. Treating refers to any process designed to separate crude oil from water and foreign contaminates carried along with it from the reservoir. Emulsion treating processes require some combination of the following: chemical addition, settling time, heat, and electrostatic coalescing.

Figure 4-35. Horizontal oil dehydrator.

4.4.2: Chemical Addition
The purpose of treating chemicals is to induce coalescence so that the oil and water will separate rapidly. Surface-active agents are absorbed at the oil–water interface, rupture the tough film (skin) surrounding the water droplets, and/or displace the emulsifying agent and force the emulsifying agent back into the oil phase. There is not a universal chemical able to break all emulsions in different crude oils. Determining the correct chemical to use is commonly done by a chemical sales representative using a bottle test (discussed earlier in this chapter).

4.4.2.1: Amount of Chemical
The amount of chemical required cannot be predicted accurately from bottle tests. The only reliable method of determining the amount of chemical to use is to run tests in the field. When changing to a new chemical or starting up a new treating system, one must first use an excess (0.25 gallon per 100 barrels, approx. 60 ppm ) of chemical and then gradually reduce the amount to the minimum amount that will produce the desired results.
When determining the amount of chemical to add, one must make certain no other changes are being made in the facility. Temperature should remain constant during the test; otherwise, it is impossible to determine which change, chemical or temperature, has caused a certain effect.
The amount of chemical added can vary from 1 gallon per 400 barrels to 0.5 gallon per 1000 barrels (60 to 10 ppm.). Concentrations higher than 120 ppm should be investigated for possible errors such as incorrect chemical being used or the method of chemical addition being wrong.
Too much chemical can be the cause of a very tight emulsion that will not break down.
Chemicals should be added continuously as possible during the entire production period and at a rate related to the production rate. Even though some residual chemical is held in the treater or gunbarrel, chemicals cannot be batched and be expected to do an adequate treatment. Chemicals cannot act properly unless they are thoroughly mixed with the emulsion.
The farther upstream, a minimum of 200 feet, from a treater or gunbarrel the chemicals are added the better the mixing and thus the better the treatment. The ideal location for injection is at the manifold before the fluid enters a separator. In some cases an emulsion that is difficult to treat may break quite easily if a chemical pump is set at the well. It is not uncommon for one well in a field to cause most of the trouble. Setting a pump at this well can increase efficiency and reduce the amount of chemicals required to break the emulsion.

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4.4.2.2: Bottle Test Considerations

The best demulsifier is the compound that results in the most rapid and complete separation of the phases at a minimum concentration. The important characteristics in the bottle test will be dictated by the production needs and the behavior of the system.

Water Drop-Out Rate
In high water volume systems a chemical that creates a fast water drop-out rate is necessary to make the system function as designed. When free water knockouts are used, the speed of water drop-out may become the most important factor. Chemicals with fast water drop-out characteristics are sometimes incomplete in treatment requiring other chemicals for final separation. In low water volume systems (fields with facilities having longer than normal residence times), the rate of water drop-out may be of minor significance in selecting the best demulsifier. In all cases, the rate of water drop-out should be noted and recorded.
Sludge
When oil, water and sediment collect together without breaking to separate water, oil and solid phases, the result is called sludge. Sludge is stabilized by finely divided solids and other contaminates to form pads that cause a secondary emulsion located between the oil and water. Depending upon the system and sludge stability, interface sludge may or may not cause a problem. Loose interface sludge can be detected by swirling the test bottle about its axis, and if the material is loose, it will break.
Interface
The desired interface is one that has shiny oil in contact with the water (mirror interface). The interface, when using a new chemical, should be as good as, if not better than, that formed by the chemical being replaced.
Water Turbidity
The turbidity (clarity) of the water is very difficult to interpret in the bottle test and correlate to facility behavior. Clear water is definitely the desired result.
Oil Color
Emulsions have a hazy appearance when compared to the bright color of treated oil. As a crude oil emulsion separates, the color tends to brighten.
Brightening of oil can be encouraging, but it can also be deceptive if taken as the sole qualification for chemical selection. While bright color is no guarantee of a successful chemical, lack of it assures that the compound is not worthy of further consideration.
Centrifuge Results
An important quality in the final evaluation is the centrifuge results. It is always good practice to make a centrifuge grind out to determine the final accurate amount of B. S&W entrained in the oil.

4.4.2.3: Chemical Selection
A thorough understanding of the treating equipment and its contribution to the treatment are necessary before chemical selection can be made. If little agitation is available, a fast-acting chemical is necessary. If a free-water knockout vessel is used, the water drop-out rate will be very important. If heat is unavailable, the chemical must work at ambient temperatures. Different types of vessels require different chemical actions.

Settling Tank or “Gunbarrel”
Speed is not too important since both tanks usually have a high volume to-throughput ratio. The chemical may continue acting over a relatively long period. An interface layer often develops but usually stabilizes at some acceptable thickness. An interface layer in a gunbarrel sometimes aids the treating process in that it acts as a filter for solids and unresolved emulsions. Fresh oil containing a demulsfier passing up through the interface layer helps treat the interface and prevents an excessive build-up.

Vertical Heater-Treater
The speed of chemical action is important since the volume-to-throughput ratio is usually lower than a gunbarrel or settling tank. With the higher throughput, it is harder to stabilize an interface layer, so more complete treatment is necessary in a shorter time period. Solids control may be important in controlling the interface.

Horizontal Heater-Treater
The speed of chemical action is important due to its high throughput. The large interface area and shallow depth require that the interface be fairly clean. Since this treater can tolerate only very little interface accumulation, the chemical treatment must be complete. Since solids tend to collect at the interface, the chemical must also effectively de-oil any solids so that they may settle out by gravity.

4.4.2.4: Settling Time
Following the addition of treating chemicals, settling time is required to promote gravity settling of the coalescing water droplets. Figure 4-36 illustrates the effects of time on coalescence.
Emulsion treating equipment designed to provide sufficient time for settling of desired water droplet size, whatever it is free water (as in three phase separators, free-water knockouts, heater-treaters, and gunbarrels) or emulsified water(as in heater treaters and electrostatic heater treater). The time necessary for water to settle is affected by differential density of the oil and water, viscosity of the oil, size of the water droplets, and relative stability of the emulsion.

4.4.2.5: Coalescence
The process of coalescence in oil treating systems is time-dependent.
In dispersions of two immiscible liquids, immediate coalescence seldom occurs when two droplets collide. If the droplet pair is exposed to turbulent pressure fluctuations, and the kinetic energy of the oscillations induced in the coalescing droplet pair is larger than the energy of adhesion between them, the contact will be broken before coalescence is completed.
Experiments shows that:
• A doubling of residence time increases the maximum size drop grown in a gravity settler less than 19%.
For this reason, after an initial short coalescence period, adding additional retention time is not very effective for making the oil easier to treat. Very often engineers will attribute improved performance in large gunbarrel tanks to retention time when it is really due to slowing the oil velocity. This allows smaller droplets of water to separate in accordance with Stokes’ law.
• The more dilute the dispersed phase, the greater the residence time needed to “grow” a given particle size will be. That is, coalescence occurs more rapidly in concentrated dispersions. This is the reason that oil is “water-washed” by entering the treating vessel below the oil–water interface in most gunbarrels and treaters. Flocculation and coalescence therefore occur most effectively at the interface zone between oil and water, (Refer to figure 4-37).

Figure 4-36. Effect of time on coalescence. Top: emulsion without chemicals. Bottom: emulsion with demulsifier added.

Figure. 4-37. Coalescence phenomenon.

4.4.2.6: Viscosity
The viscosity of the oil continuous phase is extremely important in sizing a treater. Stokes’ law, used to determine the settling velocity of a water droplet settling through the continuous oil phase, includes the oil viscosity. As the oil viscosity increases, the settling velocity of a given droplet decreases. This requires that the treater size be increased. The oil viscosity also affects coalescence of the water droplets. As the oil viscosity increases, there is more resistance to random motion of the water droplets. Therefore, the droplets do not move as fast or as far. This decreases the energy and the frequency of droplet collisions. Thus, it is more difficult to grow large water droplets in the vessel. As the oil viscosity increases, it is also more difficult to shear the oil droplets that coalesce in the piping leading to the vessel and in the water-wash section of the vessel. The net effect is that increasing the oil viscosity increases the size of the minimum water droplet that must be removed.
By far the best situation is to have oil viscosity versus temperature data for a particular oil to be treated. Alternately, data from other wells in the same field can usually be used without significant error. This viscosity versus temperature data may be plotted on special ASTM graph paper. Such plots are usually straight lines, unless the oil has a high cloud point. If the plot is a straight line, the viscosity may then be predicted at any temperature.
Laboratory testing of a particular oil at various temperatures is the most reliable method of determining how an oil behaves. ASTM D 341 outlines a procedure where the viscosity is measured at two different temperatures and then, either through a computation or on special graph paper, the viscosity at any other temperature can be obtained.
As a rule, with crude of 300API and higher, the viscosity is low and not important is separation process. Between 300API and 110API, the viscosity becomes more important, until in some cases it is impossible to process very low gravity crudes without a diluent to reduce the viscosity. The use of a diluent is not unusual for crude oil below 140API.
With virtually any crude oil the viscosity change with temperatures can be an excellent guide to minimum crude processing temperatures. An ASTM chart of the viscosity versus temperature is useful to detect the paraffin formation or cloud point of the crude as shown in Figure 4-38.
There are examples of 300API crude and higher that have pour points of 80 to 900F (27 to 320C).
If no data are available, the oil viscosity may be estimated by a variety of methods from the temperature and oil gravity. These methods, however, are not very accurate, as the viscosity is a function of the oil composition and not strictly the oil gravity. In fact, two oils with the same gravity at the same temperature may have viscosities that are orders of magnitude apart. In the absence of any laboratory data, Figure 4-39 may be used to estimate oil viscosities. Additional correlations that can be used to estimate crude viscosity given its gravity and temperature are discussed in Chapter 1.

4.4.2.7: Heat Effects
Adding heat to the incoming oil–water stream is the traditional method of separating the phases. The addition of heat reduces the viscosity of the oil phase, allowing more rapid settling velocities in accordance with Stokes’ law of settling. For some emulsifying agents, such as paraffins and asphaltenes, the addition of heat deactivates, or dissolves the emulsifier and thus increases its solubility in the oil phase. Treating temperatures normally range from 100–1600F (38–700C). In treating of heavy crudes the temperature may be as high as 3000F (1500C).
Adding heat can cause a significant loss of the lower-boiling-point hydrocarbons (light ends). This results in “shrinkage” of the oil, or loss of volume. The molecules leaving the oil phase may be used as fuel, vented, or compressed and sold with the gas. Even if they are sold with the gas, there will probably be a net loss in income realized by converting liquid volume into gas volume.
Figure 4-40 shows the amount of shrinkage that may be expected from a typical 330API gravity crude oil.
Increasing the temperature at which treating occurs has the disadvantage of making the crude oil that is recovered in the storage tank heavier and thus decreasing its value. Because the light ends are boiled off, the remaining liquid has a lower API gravity. Figure 4-41 shows the API gravity loss for a typical crude oil.
Increasing the temperature may lower the specific gravity, at the treater operating pressure, of both the oil to be treated and the water that must be separated from it. However, depending on the properties of the crude, it may either increase or decrease the difference in specific gravity. In most cases, if the treating temperature is less than 2000F (930C), the difference between the oil and water specific gravities (ΔSG) is constant and thus changes can be neglected.

Finally, it takes fuel to provide heat, and the cost of fuel must be considered.
Thus, while heat may be needed to treat the crude adequately, the less heat that is used, the better.

The gas liberated when crude oil is heated may create a problem in the treating equipment if the equipment is not properly designed. In vertical heater-treaters and gunbarrels the gas rises through the coalescing section.
If much gas is liberated, it can create enough turbulence and disturbance to inhibit coalescence. The usual oil-field horizontal heater-treater tends to overcome the gas liberation problem by coming to equilibrium in the heating section before introducing the emulsion to the settling-coalescing section.
If properly and prudently done, heating an emulsion can greatly benefit water separation. However, if a satisfactory rate of water removal can be achieved at the minimum temperature delivered into a process, there may be no reason to suffer the economic penalties associated with adding heat.

Figure 4-38, Oil viscosity vs. gravity and temp. (Courtesy of Paragon Eng. Services, Inc.)

Figure 4-39, typical viscosity-temperature curves for crude oils. (Courtesy of ASTM D-341.)
(Light crude oil (300–400API), Intermediate crude oil (200–300), & Heavy crude oil (less than 200 API)

Figure 4-40. Percent loss by volume as a function of temperature for a 330API gravity crude oil. Left
Figure 4-41. API gravity loss as a function of temperature for a 330API gravity crude oil. Right

4.4.2.8: Electrostatic Coalescers
Coalescing of the small water drops dispersed in the crude can be accomplished by subjecting the water-in-oil emulsion to a high-voltage electrical field. When a non-conductive liquid (oil) containing a dispersed conductive liquid (water) is subjected to an electrostatic field, the conductive particles or droplets are caused to combine by one of three physical phenomena:
• The droplets become polarized and tend to align themselves with the lines of electric force. In so doing, the positive and negative poles of the droplets are brought adjacent to each other. Electrical attraction brings the droplets together and causes them to coalesce.
• In an A-C field, due to inertia, small droplets vibrate over a larger distance than larger droplets promoting coalescence. In a D-C field the droplets tend to collect on the electrodes, forming larger and larger drops until eventually they fall by gravity.
• The electric field tends to distort and thus weaken the film of the emulsifier surrounding the water droplets. Water droplets dispersed in oil and subjected to a sinusoidal alternating-current field will be elongated along the lines of force during the first half cycle. As they are relaxed during the low-voltage portion, the surface tension will pull the droplets back toward the spherical shape. The same effect is obtained in the next half of the alternating cycle. The weakened film is thus more easily broken when droplets collide, making coalescence more likely.
Whatever the actual mechanism, the electric field causes the droplets to move about rapidly in random directions, which greatly increases the chances of collision with another droplet. When droplets collide with the proper velocity, coalescence occurs.
The greater the voltage gradient is, the greater the forces causing coalescence will be. However, experimental data show that at some gradient the water droplet can be pulled apart and a strong emulsion can be developed. For this reason, electrostatic treaters are normally equipped with a mechanism for adjusting the gradient in the field (Refer to figures 4-42, 4-43, and 4-44).

4.4.2.9: Water Droplet Size and Retention Time
The droplet diameter is the most important single parameter to control to aid in water settling since this term is squared in Stokes’ law’s settling equation. A small increase in diameter will create a much larger increase in settling velocity. Thus, in sizing treating equipment, it is necessary to predict a droplet diameter, which must be separated from the oil to meet a desired B.S&W specification.
It would be extremely rare to have laboratory data of droplet coalescence for a given system. Qualitatively, we would expect droplet size to increase with retention time in the coalescing section and with heat input, which excites the system, leading to more collisions of small droplets.
Droplet size could be expected to decrease with oil viscosity, which inhibits the movement of the particles and decreases the force of the collision.

Except for providing some minimal time for initial coalescence to occur, increasing retention time in a crude oil treating system may not be very cost-effective. Consequently, in most systems one would not expect retention time to have a significant impact on increasing the water droplet diameter.
Typically, retention times vary from 10 to 30 minutes, but values outside this range are also common.

Figure. 4-42.Behavior of water droplet in DC field in electrostatic coalescing.

Figure 4-43. Dual polarity DC/AC fields in electrostatic coalescing.

Figure 4-44. Effect of AC current on droplets.

4.5: Heat Required
The heat input and thus the fuel required for treating depend on the temperature rise, amount of water in the oil, and flow rate. Heating water requires about twice as much energy as it does to heat oil. For this reason, it is beneficial to separate any free water from the emulsion to be treated with either a free-water knockout located upstream of the treater or an inlet free-water knockout system in the treater itself.
Adding one Btu of heat to one lbm of water will raise the temperature of water by 1°F. This implies that the specific heat (cp) of water is one Btu/lbm-°F. While specific heat (cp) of oil is approximately 0.5 Btu/lbm-°F ( Adding 1 Btu to one lbm of oil will raise the temperature by 2 0F).

4.5.1: Heat duty
The heat duty is determined by the sum of the heat requirements for the oil and water as given by the following equation:
Heat required for emulsion

Qr = 15 W (ΔT) [co ρo (1-X) + cw ρw (X) ] Eq. 4-1

where
Qr = heat required, Btu/hr
W = flow rate of emulsion, bbl/d
Co = specific heat of the oil, btu/(lb-F) ( = 0.52)
Cw = specific heat of the water, btu/(lb-F) ( = 1.0)
X = volume fraction of water, fraction (0.0 to 1.0)
ρo = specific gravity of oil, water = 1.0
ρw = specific gravity of water, water = 1.0
ΔT = Temperature difference between inlet and outlet crude, 0F

Substituting values for a 35 API oil with a specific gravity of 0.8498 and a specific heat of 0.52 BTU/lb-F along with values for water of a specific gravity of 1.0 and a specific heat of 1.0 BTU/lb-F gives the following simplified equation.

Qr = W ΔT [6.44 + 8.14 (X)] Eq. 4-2

It must be remembered that the heat required is the heat delivered to the fluid and does not include any heat loss or the additional heat required for combustion efficiency.

4.5.2: Heat Loss
In determining the total heat input required for treating systems, the maximum amount of heat loss from the shell of the treating vessels or heat generating equipment should be taken into account. The heat loss for uninsulated vessels may be approximated from the following formula:
Ql = KDL (T2-Ta) Eq. 4-3

Where:
Ql = heat loss, Btu/hr.
K = Constant
= 15.7 for 20 mile/hour wind
= 13.2 for 10 mile/hour wind
= 9.8 for 5 m/hour wind
= 9.3 for still air
D = diameter of treater, ft.
L = height or length of shell, ft
T2 = treating temperature, 0F
Ta = design minimum outside ambient temperature,0F

For insulated vessels the heat loss may be estimated in the range of 5-10% of what the bare vessel heat loss would be.

The Total heat transfer (heat required), Btu/hr
Qt = Ql + Qr Eq. 4-4
4.5.3: Fire Tube Heat Flux
This term is commonly applied to the average heat transfer rate through the firetube, expressed as BTU/hour/square foot of cross sectional area.
The average firetube heat flux (Btu/hr/sq.ft. of exposed area) should be within range of 10,000 to 12,000 for glycol/water bath. The heat flux may be increased for fresh water bath application.
For oil water emulsion it is usually considered 10,000 Btu/hr/sq as a maximum value.

Example 4-1: Fire tube having 25.0 square feet of firetube surface, and rated @ 250,000 Btu/hr.
Average heat flux = 250,000/25 = 10,000 = 10,000 Btu/hr.sq.ft.

4.5.4: Firetube Heat Density
Heat released through the cross-sectional area of the fietube is regulated by the burner mixer and burner nozzle. Treaters conforming to this specification will have a maximum heat density of 15,000 BTU/hr/sq. in. for natural draft burners.

4.6: Treater Equipment Sizing
4.6.1: General Considerations
The major factors controlling the sizing of emulsion treating equipment are
• Heat input required,
• Gravity separation considerations,
• Settling equations,
• Retention time equations,
• Water droplet size.

4.6.1.1: Gravity Separation Considerations
Most oil-treating equipment relies on gravity to separate water droplets from the oil continuous phase, because water droplets are heavier than the oil. However, gravity is resisted by a drag force caused by the droplets’ downward movement through the oil. When the two forces are equal, a constant velocity is reached, which can be computed from Stokes’ law as (Chapter 2).
Vt = 1.78 x 10-6 (ΔSG) d2m / µ Eq. 4-5

Vt =terminal (settling velocity) of the droplet, ft/s,
dm =droplet diameter, microns,
µ =viscosity of the gas, cp.
ΔSG = difference in specific gravity between oil and water (water =1)

Several conclusions can be drawn from Stokes’ law:
• The larger the size of a water droplet, the larger the square of its diameter and, thus, the greater its downward velocity will be. That is, the bigger the droplet size, the less time it takes for the droplet to settle to the bottom of the vessel and thus the easier it is to treat the oil.
• The greater the difference in density between the water droplet and the oil phase, the greater the downward velocity will be. That is, the lighter the crude, the easier it is to treat the oil. If the crude gravity is 10 0API and the water is fresh, the settling velocity is zero, as there is no gravity difference.
• The higher the temperature, the lower the viscosity of the oil and, thus, the greater the downward velocity will be. That is, it is easier to treat the oil at high temperatures than at low temperatures (assuming a small effect on gravity difference due to increased temperature).

4.6.1.2: Settling Equations
The specific gravity difference between the dispersed water droplets and the oil should result in the water “sinking” to the bottom of the treatment vessel.
Since the oil continuous phase is flowing vertically upward in both vertical and horizontal treaters previously described, the downward velocity of the water droplet must be sufficient to overcome the velocity of the oil traveling upward through the treater. By setting the oil velocity equal to the water settling velocity, the following general sizing equations may be derived:

Horizontal Vessels:

dLeff = 438 FQo µo / (ΔSG) d2m Eq. 4-6

where
d = minimum vessel internal diameter, in.
Qo = oil flow rate, BOPD,
µo = oil viscosity, cp,
Leff = length of coalescing section, ft,
ΔSG = difference in specific gravity between oil and water (relative to water),
dm = diameter of water droplet, microns,
F = short-circuiting factor

If the treater has a spreader and a collector, then the spreader/collector short-circuiting factor is 1. If the treater lacks the spreader, collector, or both, then “F” should be some value greater than 1.

Vertical Vessels:

d = 81.8 [FQo µo / (ΔSG) d2m ]0.5 Eq. 4-7

Note that the height of the coalescing section for a vertical treater does not enter into the settling equation. The cross-sectional area of flow for the upward velocity of the oil is a function of the diameter of the vessel alone. This is a limiting factor in the capacity of vertical treaters.
In a horizontal vessel, the cross-sectional area for flow for the upward velocity of the oil is a function of the diameter times the length of the coalescing section.

Gunbarrels
The equations for gunbarrels are similar to those for vertical treaters since the flow pattern and geometry are the same. However, gunbarrel tanks experience a great deal of short-circuiting due to uneven flow distribution.
This is a result of the large tank diameter. The sizing equation for gunbarrels includes a short-circuiting factor “F.” This factor accounts for imperfect liquid distribution across the entire cross section of the treating vessel or tank and is a function of the flow conditions in the vessel. The larger the retention time, the larger the short-circuiting factor will be.

4.6.1.3: Retention Time Equations
The oil must be held at temperature for a specific period of time to enable de-emulsifying the water-in-oil emulsion. This information is best determined in the laboratory but, in the absence of such data, 20 to 30 minutes is a good starting point.
The retention time in the coalescing-settling section of a treater is the volume of the coalescing-settling section divided by the oil flow rate.
The volume of the coalescing-settling section is a function of the square of the vessel diameter and the length of the flow path of the coalescing section.
Depending on the specific properties of the stream to be treated, the geometry required to provide a certain retention time may be larger or smaller than the geometry required to satisfy the settling equation. The geometry of the vessel is determined by the larger of the two criteria.
The equations for retention time are as follows.

Horizontal Vessels:
d2 Leff = (tr)o Qo / 1.05 Eq. 4-8

Vertical Vessels:
d2 h = (tr)o Qo / 0.12 Eq. 4-9

Part of the overall vessel height is required to provide for water retention.
The removal of oil from the water is not a primary concern. Equations can be derived for water retention similar to the equations for oil retention. Assuming that a short-circuiting factor is not critical, the height required for water retention can be derived.

Gunbarrels
d2 h = F (tr)o Qo / 0.12 Eq. 4-10
tr = retention time, min,
Qo = oil flow, BOPD,
h = height of the coalescing section, in.,
F = short-circuiting factor

4.6.1.4: Water Droplet Size
In order to develop a treater design procedure, the water droplet size to be used in the settling equation to achieve a given outlet water cut must be determined. It would be extremely rare to have laboratory data of the droplet size distribution for a given emulsion as it enters the coalescing section of the treater.
We have seen that, after an initial period, increasing the retention time has a small impact on the rate of growth of particles. Thus, for practically sized treaters with retention times of 10 to 30 minutes, retention time would not be expected to be a determinant variable. Intuitively, one would expect viscosity to have a much greater effect on coalescence than temperature.
Assuming that the minimum required size of droplets that must be settled is a function only of oil viscosity, equations have been developed correlating this droplet size and oil viscosity.
The calculated droplet sizes were correlated with oil viscosity, and the following equations resulted:
dmi = 200 µo0.25 Eq. 4-11

, µo < 80cp,
where
dmi = diameter of water droplet to be settled from the oil to achieve 1% water cut, microns,
µo = viscosity of the oil phase, cp.

dmi = 200 µo0.4 Eq. 4-12

3<µo < 80cp

For viscosities below 3 cp, Eq.4-11 should be used.

4.6.2: Design Procedure
In specifying the size of a treater, it is necessary to determine the diameter (d), length or height of the coalescing section (Leff or h), and treating temperature or fire-tube rating. As we have seen, these variables are interdependent, and it is not possible to arrive at a unique solution for each. The design engineer must trade the cost of increased geometry against the savings from reducing the treating temperature.
The equations previously presented provide tools for arriving at this trade-off. However, because of the empirical nature of some of the underlying assumptions, engineering judgment must be utilized in selecting the size of treater to use.

4.6.2.1: Design Procedure for Horizontal Heater-Treaters
The following procedure is mostly aimed at determining the minimum size of the coalescing/settling section of the treater and the rating of the burner.
Such information will be very useful in preparing equipment specifications for vendors and for evaluating the quotations received from the vendors.
The vendors would provide the detailed design and dimensions of the treater.
1. The first step is to decide on a treating temperature. This is best determined from laboratory tests. The optimum treating temperature must provide a minimum loss of oil volume and quality along with a practical treater size. If laboratory data are not available, the treating temperature may be determined based on experience. In such cases, however, the design (following steps) may be executed for different assumed treating temperature and a final decision is made based on analysis of the design results.
3. Determine the diameter of the water droplet that must be removed, from Eq. 4-11 or 4-12.
3. Use Eq. 4-6 to obtain the relation between D and L that satisfies the settling constraint. Assume various values of D and determine the corresponding values of L from this relation.
4. Use Eq. 4-8 to obtain another relation between D and L that satisfies the retention time constraint. For the same values of D assumed in step 3, determine corresponding values of L from this relation.
5. Compare the results obtained from the above two steps and select a combination of D and L that satisfies both settling and retention time constraints.
6. For various standard diameters, develop a table of effective lengths versus standard diameters.
7. Select a treater dimensions, which satisfies the larger effective length requirements for the selected diameter.
8. Determine the heat input required using eqs. 4-2, 4-3, and 4-4.
9. Choose the nearest suitable dimensions from manufactures and vendor supply product tables.

4.6.2.2: Design Procedure for Vertical Heater-Treaters and Gunbarrels
Similar to horizontal treaters, the following procedure is primarily aimed at determining the minimum size of the coalescing/settling section of the treater and the rating of the burner.
1. Determine the optimum treating temperature that provides the minimum lose of oil volume and quality along with a practical treater size. If this is not available, the design (following steps) may be executed for different assumed treating temperature and a final decision is made based on analysis of the design results.
2. Determine the diameter of the water droplet that must be removed, from Eq. 4-11 or 4-12.
3. Use Eq. 4-7 to obtain the minimum treater diameter D that satisfied the settling constraint.
4. Repeat the above steps for different assumed treating temperatures and determine the values of D for each treating temperature.
5. Use Eq. 4-9 or 4-10 to obtain a relation between D and H that satisfies the retention time constraint. Then, assume different values of D and determine corresponding value of H from this relation.
6. Analyze the results to determine the combinations of D and H, for each treating temperature, that satisfy both settling and retention time constraints.
7. Select the treater dimensions, which satisfies the larger height requirements for the selected diameter.
8. Use Eq. 4-2 or 4-3 or 4-4, to determine the heat requirement for the selected treating temperature.
9. Choose the nearest suitable dimensions from manufactures and vendor supply product tables.

Example 4-2. Horizontal heater sizing.
Determine the heat requirement and the size of the settling/coalescing section of a horizontal heater treater for the following conditions:
Oil flow rate: 7000BPD
Inlet B.S.&W.: 15%
Outlet B.S.&W.: 1%
Oil specific gravity: 0.86
Oil viscosity: 45 cP at 850F
20 cP at 1050F
10 cP at 1250F
Water specific gravity: 1.06
Specific heat of oil: 0.5 Btu/lb 0F
Specific heat of water: 1.1 Btu/lb 0F
Inlet temperature: 850F
Retention time: 20 min
Treating temperature: Examine 1050F, 1250F, and no heating
Solution
Use Eq. (4-11) to determine the water droplet diameter for each treating temperature:

For T = 125 0F: dm = 200 (սo)0.25 = 200 (10)0.25 = 356 micron

For T = 105 0F: dm = 200 (20)0.25 = 423 micron

For T = 85 0F: dm = 200 (45)0.25 = 518 micron

Ignoring the effect of temperature on specific gravity, use Eq. 4-6 to determine the settling constraint for each treating temperature, and consider F = 1:
dLeff = 438 FQo µo / (ΔSG) d2m Eq. 4-6

For T = 125 0F
dLeff = 438 x 7000 x 10 / (0.2) (356)2
dLeff = 1204 in. ft (E1)

For T = 105 0F
dLeff = 438 x 7000 x 20 / (0.2) (423)2
dLeff = 1706 in. ft (E2)

For T = 85 0F
dLeff = 438 x 7000 x 45 / (0.2) (518)2
dLeff = 2559 in. ft (E3)

Use Eq. 4-8 to determine the relationship for retention time constraints:
d2 Leff = (tr)o Qo / 1.05 Eq. 4-8

d2 Leff = 20 x 7000 / 1.05 = 133,333 in2 ft (E4)

Assume different values for D and determine the corresponding values of L from Eqs. (E1)–(E4). The results are summarized in the following table and are plotted for comparison.

Table. 4-1 Solution of example 4-2.

Analyzing the tabulated/plotted results yields the following conclusions:
1. Any combination of D and L that exists in the plot area below the retention time curve is not acceptable. [Value of length ( eqs. E1, E2, and E3) must be bigger than the retention time value (eq. E4)]
2. For the treater diameters selected in the table, only the values of L shown in bold are acceptable, as they satisfy both settling and retention time constraints.
3. As the treating temperature increases, the size of the coalescing/settling section decreases.
4. There is no need to treat the emulsion at 1250F, as the reduction in treater size is not significant, and the increased temperature would negatively affect the volume and quality of the treated oil.
5. There is a good potential of treating this oil without any heating aid, as the treater size required seems to be practical.
6. A practical and economical selection would be an 84-in.-diameter by 21-ft-long coalescing section with a burner that can provide a treating temperature of 1050F.

Fig. 4-45. Solution of example 4-2.

Now use Eq. 4-1 to calculate the heat requirement, assuming 10% heat losses:

Qr = 15 W (ΔT) [co ρo (1-X) + cw ρw (X) ] Eq. 4-1

The value will be multiplied by 1/(1-0.1) for 10% heat losses.

Qr = (1/1-0.1)15x 7000 (20) [0.5x0.86 (1-0.15) + 1.1x 1.06 (0.15)]
Qr = 1.3 MMBTU/hr

Example 4-3. Vertical heater sizing.
Determine the heat requirement and the size of the settling/coalescing section of a single-well vertical heater treater for the given data
Oil gravity = 40 0API, 0.875 SG
Oil flow rate = 2,000 BOPD
Inlet oil temperature= 90 0F
Water SG = 1.04
Inlet BS&W = 10%
Outlet BS&W = 1%
co =0.5
co = 1.0
Retention time: 20 min
Assuming droplet diameters 325, 301, and 270 for treating temperature 90, 100, and 120 0F respectively.
Crude oil viscosity 7,5.1 ,3.3 cP. At given temperature respectively.

Use the next table which include the difference in specific gravity at different treating temperatures.

Table. 4-2. Example 4-3.

Solution
Use Eq. 4-7 to determine the minimum diameter at the three treating temperatures:
d = 81.8 [FQo µo / (ΔSG) d2m ]0.5 Eq. 4-7

For T = 90 0F
d = 81.8 [2000 x 7 / (0.215) (325)2 ]0.5
d = 64 in

For T = 100 0F
d = 81.8 [2000 x 5.1 / (0.215) (301)2 ]0.5
d = 59 in

For T = 120 0F
d = 81.8 [2000 x 3.3 / (0.215) (270)2 ]0.5
d = 53

Table. 4-3 Solution of example 4-3.

Now, use Eq. 4-9, for the retention time constraint:
d2 h = (tr)o Qo / 0.12 Eq. 4-9

d2 h = 20 x 2000 /0.12 = 333,3 in3
Assume different values for D and determine corresponding values of H from the above relation. The results are plotted as follows:

Fig. 4-46. Solution of example 4-3.

Table. 4-4 Solution of example 4-3.

From the figure, all diameters and heights that fall below the retention time curve are not acceptable. For the three treating temperatures, a coalescing section height equal to the value at the intersection with the retention time curve, or larger, will satisfy both retention time and settling constraints.

The burner rating in determined from
Qr = 15 W (ΔT) [co ρo (1-X) + cw ρw (X) ] Eq. 4-1

For 120 0F = 0.45 MMBBTU/hr
For 100 0F = 0.15 MMBBTU/hr
For 90 0F = 0 MMBBTU/hr

4.7: Practical Considerations
Successful treatment of emulsions, depending on specific emulsion characteristics, can be treated by low temperature with or without adding chemicals, or chemicals with or without heat. Some fields having high water cut (e.g., 95%) can be treated successfully without heat or chemicals, but require extremely long retention times. It is better to use chemicals instead of heat from the standpoints of installation, maintenance, and operating costs. The following discussion provides some general guidelines to help one select the right oil treating equipment configuration for a specific application.

4.7.1: Gunbarrels with Internal/External Gas Boot
Gunbarrels (wash tank with internal/external gas boot) should be considered when isolated, high-salt-water percentage production is indicated, provided retention time requirements do not make gunbarrel sizing impractical. When used without heat, the vessel should provide ample settling time, e.g., 12 to 24 hr.
Sufficient retention time allows some storage of basic sediment during cold weather when chemical efficiency declines.
The basic settlement is cleaned from the tank during warm weather and by periodically rolling (circulating) the gunbarrel.

4.7.2: Heater-Treaters
A heater-treater should be considered in fields requiring heat to break the emulsion. Good practice is to install a slightly larger (+10%) heater-treater than is necessary. This allows extra capacity for unforeseeable production increases (normally water), reduction in the amounts of treating chemical used, and startup of a cold unit. A reduction in chemical cost can easily pay for the additional cost of a larger treater in a few years. Depending on the characteristics of the oil and the efficiency of the chemical, retention times range between 10 to 60 minutes.

4.7.3: Electrostatic Heater-Treaters
An electrostatic heater-treater should be considered in fields with maximum salt content specifications imposed [10 to 30 lb per thousand barrels (PTB)], any time the BS&W must be reduced below 0.5%, and offshore facilities where space and/or heat is limited.

Other configuration considerations that the designer may be required to evaluate are free-water knockout instead of a gunbarrel and using an electrostatic heater-treater instead of a heater-treater.

yasserkassem · Post by **yasserkassem** » Sun Mar 22, 2020 1:31 pm

Chapter - 5
Crude Oil Desalting

---------------
Chapter 5 185
Crude Oil Desalting 185
5.1: Introduction 185
5.1.1: Salt Content 185
5.1.2: Desalting Process 186
5.2: Equipment Description 186
5.2.1: Desalters 186
5.2.2: Mixing Equipment 186
5.3: Process Description 188
5.3.1: Single-Stage Desalting 189
5.3.2: Two-Stage Desalting 189
5.4: Electrostatic Desalting Voltage 189
5.5: Operating Parameters Effects 191
5.6: Design Consideration 191
5.7: Troubleshooting 192
----
Chapter 5

Crude Oil Desalting

5.1: Introduction
The process of removing water-soluble salts from an oil stream is called oil desalting. Nearly all crude oil contains entrained water, which always contains dissolved salts, specifically sodium chloride. The majority of the produced salt water is removed by separation and the oil treating process (dehydration). However, a small amount of entrained water remains in the crude oil.
Refineries usually specify in purchase contracts a maximum salt content, as well as maximum water content. A common salt specification would be 10 to 20 pounds per thousand barrels. To satisfy the refinery specification, upstream production facilities may be required to perform oil desalting.
This chapter describes the methods and equipment commonly used to desalt crude oil.

5.1.1: Salt Content
The amount of salt in the crude oil is a function of the amount of the brine that remains in the oil WR (% BS&W), and of its salinity SR in parts per million as sodium chloride (ppm). In other words, this relationship could be written in the following functional form:

Salt content (PTB) = 0.35 ρBrine SR WR/(100-WR) Eq. 5-1

Where
PTB = Pounds salt per 1000 barrel crude oil.
WR = (%BS&W)
SR = Salinity as sodium chloride (= 1.65 x Salinity as Cl-)
ρBrine = Density of brine.

So, Desalting can be achieved by:
1- Reducing the water content.
2- Reducing the salinity of the remnant water (reduce salt concentration of remnant water)

The method of reducing the PTB by lowering the quantity of remnant water WR is usually referred to as the treating process of oil dehydration. This was the main theme of the last chapter. The other alternative of reducing the PTB is to substantially decrease the dissolved salt content of the remnant water (i.e., its concentration, SR). This practice is the one we are dealing with in this chapter and is known as desalting.

Example 5-1: Find the PTB of a crude oil having 1% by volume remnant water if its concentration is estimated to be 40,000 ppm and brine density is 1.05 Kg/l.
Solution
Salt content = 0.35 x 1.05 x 40,000 x 1 /(100-1) = 148 PTB

5.1.2: Desalting Process
Desalting process may be carried out by:
1- Reduce the water content only, in case of the salinity of remnant water is not high.
2- Addition of small percentage of fresh water (water with low salinity) to the crude oil (called dilute water), mix the water with the crude oil, and dehydrate the crude oil. At least if the same water content achieved, the salt content will be lowered because of reduction in remnant water salinity.
5.2: Equipment Description
5.2.1: Desalters
Since the salt content is directly related to the amount of residual water, the best desalters remove as much water as possible. Any device that removes water from oil can be used as a desalter. However, the majority of desalters employed are horizontal electrostatic treaters. These treaters will produce the lowest residual water level of all treaters. Figures 5-1, 2, & 3, illustrate a conventional horizontal electrostatic treater of the type typically used in desalting operations. Because very low water contents are required, the crude is usually pumped through the desalter at pressures above its bubble point. In addition, the temperature of the crude to be desalted is determined by upstream heat exchangers or heater treater.

Figure 5-1. Horizontal electrostatic heater-treater.

5.2.2: Mixing Equipment
Globe Valves
A manual globe throttling valve is one of the simplest methods to promote the mixing of dilution water and salt water entrained in oil stream. The pressure drop resulting from forcing the oil and water through this manual valve is used to shear the water droplets and mix the droplets in the oil. The major disadvantage of any manual valve is its inability to automatically adjust for changes in oil flow rate. As the flow rate varies, the pressure drop, and thus the mixing efficiency, varies. Therefore, if the oil flow rate increases significantly, the pressure drop may increase to the point where the resulting mixed emulsion is impossible to treat.
It is possible to automate the globe valve to avoid “over mixing”. A differential pressure controller is used to control the pressure drop through the globe valve. This system automatically adjusts for changing flow rates and maintains a set pressure drop.
The pressure drop through the mixing valve varies from 10 to 50 psi. The required pressure drop can be decreased if a premixing device is installed upstream of the mixing valve.

Figure. 5-2 Desalter.

Figure. 5-3 Desalter.

Figure 5-4. Schematic of a spray nozzle system for premixing water and oil.

Spray Nozzles
Upstream premixing is commonly performed with either spray nozzles or static mixers.
As shown in Figure 5-4, one common method of premixing the water and oil involves using a system of spray nozzles. Water is pumped through the nozzles and then distributed throughout the oil stream. These systems are effective and are usually less expensive than static mixers.

Static Mixers
Static mixers use pieces of corrugated plate, as shown in Figure 5-5.
These mixers typically divide into many parallel paths which divide and recombine the fluid as the flow passes through the mixer. The alternate layers of corrugations are perpendicular to each other so that the fluid must pass through a series of relatively small openings. This mixer shears the water droplets to a much smaller size than the old mixers. These mixers produce a narrow range of droplet sizes. This is a result of two opposing phenomena. Large droplets are sheared by the mixing action in the small openings, while at the same time these mixers provide large surface areas where small droplets may collect and coalesce. Theoretically, the coalescing ability improves the performance of the dehydration equipment due to the reduction in the number of very small droplets which makes dehydration easier and decreases the chances of creating a stable, untreatable emulsion during the mixing process.

Static mixers are sized to provide an average droplet size using empirical equations based on test data. The average droplet size for desalting is roughly between 250 and 500 microns. The average droplet size is a function of the oil flow rate. The primary disadvantage of static mixers is that they may not be adjusted as the flow varies. Therefore, if the oil flow will vary over a range of 3 to 1, or more, static mixers should not be used as the only mixing device.

Figure 5-5. Static mixer.
5.3: Process Description
Most of the salt contained in crude oil is dissolved in the small water droplets. Since water is the salt carrier, removing the water will remove the salt from the crude. The salt content of the water is expressed as parts per million (ppm) equivalent sodium chloride. Salinity may range from 0 to over 150,000 ppm. Desalting is required when the amount of salt contained in the entrained water after treating is higher than some specified amount.
For example, assume a heater-treater is used for dehydration and it yields oil that is 0.5% water, each thousand barrels of dehydrated oil includes 5 bbls of water. If we next assume the water has a low salt content, say 10,000 ppm NaCl, then each barrel of water would contain approximately 3.5 pounds of salt. With 5 bbls of water per thousand barrels of oil, the oil would then contain approximately 17.5 PTB (pounds per thousand barrels). If the purchase agreement specified 10 PTB or less, some desalting, or a more efficient dehydrator, would be required.
In this example, an electrostatic treater might be all that is required to achieve an oil outlet that contains less than 0.3% water. This example assumed a low salt content. If the water had a high salt content, say 200,000 ppm NaCl, there would be approximately 70 pounds of salt per barrel of water (lb/bbl). In this case, even dehydrating to 0.1% leaves 70 PTB. To reach the required 10 PTB, desalting would be required.
The desalting process involves two steps. The first step is to mix fresh water with entrained produced water. This will lower the produced water salinity by diluting the salt. The second step is dehydration which is the removal of the water from the crude. This dilution and dehydration produces a lower salinity in the residual water in the crude oil. The dilution water in desalting does not have to be fresh. Any water with a lower salt content than the produced water can be used.

5.3.1: Single-Stage Desalting
Figure 5-6 is a schematic of a single-stage desalting system. In this system, the dilution water is injected into the oil stream and then mixed.
The oil then enters the desalter where the water is removed. To reduce dilution water requirements, the crude oil may be dehydrated prior to the desalting process, this removes the bulk of the produced water prior to desalting and increase dilution efficiency.

Figure 5-6. Schematic of a single-stage desalting system.

5.3.2: Two-Stage Desalting
Figure 5-7 is a schematic of a two-stage desalting system with dilution water recycling capability. In this system; the water removed in the second stage is pumped back to the first stage. The addition of this recycle provides for some dilution of the salt water prior to the first stage. This further reduces the dilution water requirement compared to a single-stage dehydrator and desalter system. If further desalting is needed, it is possible to add more stages in a similar manner.
5.4: Electrostatic Desalting Voltage
In desalters, an external electric field is applied to coalesce the small water droplets and thus promote settling of the water droplets out of the oil. The electric field may be applied in any of the following manners:
1. AC field devices for water-rich emulsions. Alternating current (ac) is applied, which alternates the polar water molecule arrangements leading to better coalescence. A schematic diagram of ac electrostatic coalescence is shown in Figure 5-8.

Figure 5-7. Schematic of a two-stage desalting system with a recycle stream.

Figure 5-8. Effect of AC current on droplets.

2. AC/DC field for maximum dehydration. A combination of ac and dc (direct current) is used in this case. The basic configuration of this process is shown in Figure 5-9. the ac is produced in the zone beneath the electrodes, whereas the dc field is produced between adjacent electrodes. This arrangement achieves maximum water removal.

Figure 5-10, illustrates the effect of DC field on water droplet, while figure 5-9 illustrate the effect of AC on water droplets. (More details were presented in previous chapter).

Figure 5-9. Dual polarity DC/AC fields in electrostatic coalescing.

Figure. 5-10.Behavior of water droplet in DC field in electrostatic coalescing.
5.5: Operating Parameters Effects
The efficiency of desalting is dependent on the following parameters:
1. Water–crude interface level. This level should be kept constant; any changes will change electrical field and perturbs electrical coalescence.
2. Desalting temperature. Temperature affects water droplet settling through its effect on oil viscosity; therefore, heavier crude oils require higher desalting temperatures.
3. Wash water ratio. Heavy crudes require a high wash water ratio to increase electrical coalescence.
4. Pressure drop in the mixing valve. A high-pressure-drop operation results in the formation of a fine stable emulsion and better washing. However, if the pressure drop is excessive, the emulsion might be difficult to break.
5. Type of demulsifiers. Demulsifiers are added to aid in complete electrostatic coalescence and desalting. They are quite important when heavy crudes are handled.
5.6: Design Consideration
The following major parameters are considered when designing the desalting system:
1. Number of desalting stages
2. Dehydration levels achieved
3. Salinity of the brine in the crude
4. Efficiency of valve mixing
5. Salinity of dilution water
6. Target PTB specification
5.7: Troubleshooting
Table 5:1 lists some ‘‘tips’’ that are helpful in solving some of the operating problems or troubles that are of significance to the desalting process.

Table 5-1, Troubleshooting of desalting process.

yasserkassem · Post by **yasserkassem** » Sun Mar 22, 2020 1:32 pm

Chapter- 6

Crude Oil Stabilization

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Chapter 6 193
Crude Oil Stabilization and Sweetening 193
6.1: Introduction 193
6-1-1: Crude oil treatment steps 193
6.2: Process Schemes 194
6.2.1: Multi-Stage Separation 194
6.2.2: Oil Heater-Treaters 194
6.2.3: Liquid Hydrocarbon Stabilizer 195
6.2.4: Cold-Feed Stabilizer 197
6.2.5: Stabilizer with Reflux 197
6.3: Stabilization Equipment 199
6.3.1: Stabilizer Tower 199
6.4: Stabilizer Design 205
6.5: Crude Oil Sweetening 206
6.6.1: Stage vaporization with stripping gas. 206
6.6.2: Trayed stabilization with stripping gas. 207
6.6.3: Reboiled trayed stabilization. 208

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Chapter 6

Crude Oil Stabilization and Sweetening

6.1: Introduction
Once degassed and dehydrated–desalted, crude oil is pumped to gathering facilities to be stored in storage tanks.
These liquids contain a large percentage of methane and ethane, which will flash to gas in the tank. This lowers the partial pressure of all other components in the tank and increases their tendency to flash to vapors. Stabilization is the process of increasing the amount of intermediate (C3 to C5) and heavy (C6+) components in the liquid phase. In an oil field this process is called crude stabilization and in a gas field it is called condensate stabilization.
In stabilization, adjusting the pentanes and lighter fractions retained in the stock tank liquid can change the crude oil gravity. The economic value of the crude oil is accordingly influenced by stabilization because of the following reasons:
1- Liquids can be stored and transported to the market more profitably than gas.
2- It is advantageous to minimize gas losses from light crude oil when stored.
This chapter deals with methods for stabilizing the crude oil to maximize the volume of production as well as its API gravity, against two important constraints imposed by its vapor pressure and the allowable hydrogen sulfide content.
In addition, the requirement to treat the oil at high temperature is more important than stabilizing the liquid and may require the flashing of both intermediate and heavy components to the gas stream.
Gas condensate, on the other hand, may contain a relatively high percentage of intermediate components. Thus, some sort of condensate stabilization should be considered for each gas well production facility.
The most common method used to remove the light components from hydrocarbon liquids before the liquid enters a stock tank or a pipeline is stage separation.
Crude oil is considered ‘‘sweet’’ if the dangerous acidic gases are removed from it. On the other hand, it is classified as ‘‘sour’’ if it contains as much as 0.05 ft3 of dissolved H2S in 100 gal of oil. Hydrogen sulfide gas is a poison hazard because 0.1% in air is toxically fatal in 30 min.
Additional processing is mandatory—via this dual operation—in order to release any residual associated gases along with H2S present in the crude.
A stabilizer can achieve a stable specification product with a higher liquid recovery, but usually results in higher capital expenditures’ (CAPEX) and operating expenses (OPEX). The addition of a stabilizer requires additional space which is normally not a factor for onshore applications, but may be a major consideration for an offshore installation.

6-1-1: Crude oil treatment steps
Produced hydrocarbons from wells normally flow to a separator for removal of the hydrocarbon gas. The hydrocarbon crude or condensate oil outflow from the separator usually goes through additional stages of separation or treatment before reaching the sales point. In each of these stages the liquid reaches near equilibrium at a different condition of pressure and temperature thus to some extent “stabilizing” the crude or condensate.
The following methods of crude stabilization are normally used:
• Multi-stage separation
• Weathering in a stock tank
• Heater-treater after separation
• Stabilizer.
The method one selects for stabilization depends primarily on contract specifications and economics. Factors that favor the installation of a stabilization unit include:
• An oil contract specification that requires a low crude vapor pressure that cannot easily be obtained by stage separation.
• A sour crude with a contract specification that limits the H2S content to less than 60 ppm.
• Condensate production with 500API or higher and flow rates in excess of 5,000 bpd.
6.2: Process Schemes
6.2.1: Multi-Stage Separation
Figure 6-1 shows a multi-stage separation system. This is the most common method of separating oil and gas. This system typically requires from two to four separation stages, each occurring in a separator vessel.

Figure 6-1. Schematic of a three-stage separation system.

6.2.2: Oil Heater-Treaters
Three-phase separators, which utilize gravity separation, often are not adequate to separate the water from the oil. Heating the emulsion is commonly used to break the emulsion. Heater treaters not only improve the oil-water separation process, but also stabilize the crude by vaporizing the light hydrocarbons prior to the crude flowing to an atmosphere pressure storage tank. Utilizing heater-treaters alone often results in higher than desired losses of intermediate components to the vapor phase when the hot crude is flashed entering the storage tank.
The crude departing the treater can be cooled before going to the storage tank by exchanging heat with the colder emulsion upstream of the treater. This will lead to fewer vapor losses and will help stabilize the intermediate components when the crude is flashed at storage tank conditions. For small flow rates, the oil-treating temperature is kept as low as possible to prevent stock tank losses, since the treated oil will normally go directly to the stock tank without cooling.

6.2.3: Liquid Hydrocarbon Stabilizer
It is possible to stabilize a hydrocarbon liquid at constant pressure by successively flashing the hydrocarbon liquid at increasing temperatures as shown in Figure 6-2. At each successive stage the partial pressure of the intermediate components is higher than it could have been at that temperature if some of the lighter components had not been removed by the previous stage. It would be very costly to arrange a process as shown in Figure 6-2 and thus never done. Instead, the same effect can be obtained in a tall, vertical pressure vessel with a cold temperature at the top and a hot temperature at the bottom. This unit is called a “stabilizer.”

Figure 6-2. Multiple flashes at constant pressure and increasing temperature.

A stabilizer applies the same principles as multi-stage separation except that the flashes take place in a stabilizer tower operating at a constant pressure, but with varying temperatures. The stabilizer tower is normally a trayed vertical pressure vessel; however, structured packing may also be used. As heat is added to the bottom of the stabilizer tower, vapors are generated on the bottom tray. The hot vapors rise to the tray above, where they bubble through the liquid. The liquid is heated by the hot vapors, which vaporize some of the hydrocarbon liquid. The vapors, in turn, are cooled by the liquid, and a portion of the vapor is condensed.
This process of vaporization and condensation is repeated on each tray in the stabilizer tower. As the liquids fall down the stabilizer tower, the heavier hydrocarbons are condensed so that the hydrocarbon liquids leaving the stabilizer tower contain almost none of the light hydrocarbon components, and the vapor leaving the top of the stabilizer tower contains almost none of the heavier components.
The vapor pressure of the liquid hydrocarbon leaving the bottom of the tower is controlled by controlling the stabilizer tower pressure and bottom temperature. At a constant pressure, the liquid hydrocarbon product’s vapor pressure can be increased by lowering the bottom temperature, or decreased by increasing the bottom temperature.

Figure 6-3 illustrates a liquid hydrocarbon stabilizer system. The well stream flows to a high pressure, three-phase separator. Liquids containing a high fraction of light ends are cooled and enter the stabilizer tower at a pressure between 100 to 200 psi.

Figure 6-3. Cold-feed stabilization system.

As the hydrocarbon liquid falls from tray to tray in the stabilizer tower, it is heated by the hot gases bubbling through the liquid. On each tray some of the liquids are vaporized and some of the hot gases are condensed. The liquids falling down the stabilizer tower become richer and richer in heavy hydrocarbon components and leaner and leaner in light hydrocarbons. At the bottom of the stabilizer tower, some of the liquids are cycled to a reboiler where they receive heat to provide the necessary bottom temperature which is normally in the range of 2000 to 4000F. The reboiler could be a direct-fired bath, an indirect-fired bath, or a heating media exchanger. For a specific bottom product’s vapor pressure, a lower stabilizer tower operating pressure requires a lower bottom temperature, but more compression is required for the overhead vapors.
The hydrocarbon liquid leaving the stabilizer tower at the bottom tray temperature is in equilibrium with the vapors and is at its bubble point.
The liquid leaving the stabilizer tower is cooled before going to storage or pipeline. The hydrocarbon vapors leaving the top of the stabilizer tower are in equilibrium with the liquids on the top tray and are at their dew point.
One design consideration that needs to be addressed in the design of a stabilizer system is whether to use a cold feed or reflux. A cold-feed stabilizer without reflux such as that shown in Figure 6-3 does not achieve as good a split between the light and heavy components as a column with reflux (see Figure 6-4 and the following discussion); thus, recoveries are not as high. However, a stabilizer with reflux requires additional equipment, higher CAPEX, and higher OPEX, but achieves a higher recovery. Descriptions of both a cold-feed stabilizer and a stabilizer with reflux follow.

6.2.4: Cold-Feed Stabilizer
A conventional stabilizer tower is a distillation column with a reboiler, but no overhead condenser (refer to Figure 6-3). The lack of an overhead condenser means that there is no liquid reflux from the overhead stream.
Thus, the feed is introduced on the top tray and must provide all the cold liquid for the stabilization tower. Since the feed is introduced on the top tray, it is important to minimize the flashing of the feed so that intermediate components are not lost overhead. To lower the feed stream temperature and reduce flashing, a cooler is sometimes added on the inlet feed stream.
Adding a cooler on the inlet feed stream lowers the temperature of the inlet hydrocarbon liquid, lowers the fraction of intermediate components that flash to vapor on the top tray and increases the recovery of these components in the liquid bottoms. However, the colder the feed, the more heat is required from the reboiler to remove light components from the liquid bottoms. If too many light components remain in the liquid, the vapor pressure limitations for the liquid may be exceeded. Light components may also encourage flashing of intermediate components (by lowering their partial pressure) in the storage tank. There is a balance between the amount of inlet cooling and the amount of reboiling required.
The hydrocarbon liquid out the bottom of the stabilizer tower must meet a specified vapor pressure. The stabilizer tower is designed to maximize the molecules of intermediate components in the liquid without exceeding the vapor pressure specification. This is accomplished by driving the maximum number of molecules of methane and ethane out of the liquid and keeping as much of the heavier ends as possible from going out with the gas. The hot liquid from the stabilizer is at its bubble point at the pressure and temperature in the stabilizer. It must be cooled sufficiently to avoid flashing when it enters the atmospheric storage tank.

The overhead gas can be used as fuel, or compressed and included with the sales gas. Any water that enters the column in the feed stream will collect in the middle of the column due to the range of temperatures involved. This water cannot leave with the bottom product or with the overhead stream; therefore, provisions should be made to remove this water from a tray near the middle of the column. The heating of the liquid hydrocarbon in the stabilizer tower acts as a demulsifier to remove water from hydrocarbon liquid. The excellent water-separating ability of the stabilizer usually eliminates the need for a hydrocarbon liquid dehydration system.

6.2.5: Stabilizer with Reflux
Figure 6-4 shows a typical stabilizer system with reflux and a feed/bottom heat exchanger. In this configuration, the well fluid is heated by the bottom product and injected into the stabilizer tower, below the top, where the temperature in the stabilizer tower is equal to the temperature of the feed. The stabilizer tower’s top temperature is controlled by cooling and condensing part of the hydrocarbon vapors leaving the stabilizer and pumping the resulting hydrocarbon liquids back to the tower. This replaces the cold feed configuration and allows better control of the overhead product and, consequently, slightly higher recovery of the heavier components. This configuration minimizes the amount of flashing.
The principles of this configuration are the same as in a cold-feed stabilizer or any other stabilizer tower. As the liquid falls through the tower, it goes from tray to tray, and gets increasingly richer in the heavier components and increasingly leaner in the lighter components. The stabilized hydrocarbon liquid is cooled in the heat exchanger by the feed stream before flowing to the stock tank or pipeline.
At the top of the stabilizer tower intermediate components going out with the gas are condensed, separated, pumped back to the stabilizer tower, and sprayed down on the top tray. This liquid is called “reflux,” and the two-phase separator that separates it from the hydrocarbon liquid from the gas is called a “reflux tank” or “reflux drum.” The reflux performs the same function as the cold feed in a cold feed stabilizer. Cold liquid hydrocarbons strip out the intermediate components from the gas as the gas rises.
The heat required at the reboiler depends upon the amount of cooling done in the condenser. The colder the condenser, the purer the product, and the larger the percentage of the intermediate components that will be recovered in the separator and kept from going out with the gas.
The hotter the bottom temperature, the greater the percentage of light components boiled out of the bottoms. The greater the percentage of light components boiled out of the bottoms liquid, the lower the vapor pressure of the bottoms liquid.
A heat balance around the stabilizer tower is part of the design. The heat leaves the stabilizer tower in the form of vapors out the top, and the liquid bottom product has to be balanced by the heat entering in the feed and the reboiler. If the stabilizer tower has a reflux, this amount of heat has to be added to the column balance.
A stabilizer tower with reflux will recover more intermediate components from the gas than a cold-feed stabilizer tower. However, it requires more equipment to purchase, install, and operate. This additional cost must be justified by the net benefit of the incremental hydrocarbon liquid recovery, less the cost of natural gas shrinkage and loss of heating value, over that obtained from a cold-feed stabilizer.

Figure 6-4. Schematic of a typical crude stabilization with reflux and feed/bottom heat exchanger.

6.3: Stabilization Equipment
6.3.1: Stabilizer Tower
The stabilizer tower is a fractionation tower using trays or packing.
Figure 6-5 shows a stabilizer tower with bubble cap trays.
Trays, structured packing, or random packing are used in the tower to promote intimate contact between the vapor and liquid phases, thereby permitting the transfer of mass and heat from one phase to the other. The feed to the stabilizer tower normally enters near the top of a cold-feed stabilizer, and at or near the tray where the stabilizer tower conditions and feed composition most nearly match the inlet feed conditions, in stabilizer towers with reflux. The liquids in the stabilizer tower fall down through the downcomer, across the tray, over the weir and into the down-comer to the next tray. The temperature on each tray increases as the liquids drop from tray to tray. Hot gases come up the stabilizer tower and bubble through the liquid on the tray above, where some of the heavier components in the gas are condensed and some of the lighter components in the liquid are vaporized. The gas gets leaner and leaner in heavy hydrocarbons as it travels up the stabilizer tower; the falling liquids get richer and richer in the heavier hydrocarbon components. The vapors leaving the top of the stabilizer tower contain a minimum amount of heavy hydrocarbons, and the liquid leaving the bottom of the tower contains a minimum of light hydrocarbons. Stabilizer columns commonly operate at pressures between 100 to 200 psig.

6.3.1.1: Trays and Packing
The more stages, the more complete the split, but the taller and more costly the tower. Most stabilizers will normally contain approximately five theoretical stages. In a refluxed tower, the section above the feed is known as the rectification section, while the section below the feed is known as the stripping section. The rectification section normally contains about two equilibrium stages above the feed, and the stripping section normally contains three equilibrium stages.

Trays
For most trays, liquid flows across an “active area” of the tray and then into a “down-comer” to the next tray below, etc. Inlet and/or outlet weirs control the liquid distribution across the tray. Vapor flows up the stabilizer tower and passes through the tray active area, bubbling up through (and thus contacting) the liquid flowing across the tray. The vapor distribution is controlled by:
• Perforations in the tray deck (sieve trays),
• Bubble caps (bubble cap trays), or
• Valves (valve trays).

Trays are generally divided into four categories:
• Sieve trays,
• Valve trays,
• Bubble cap trays, and
• High capacity/high efficiency trays.

Sieve Trays
Sieve trays are the least expensive tray option. In sieve trays, vapor flowing up through the tower contacts the liquid by passing through small perforations in the tray floor (Figure 6-6). Sieve trays rely on vapor velocity to exclude liquid from falling through the perforations in the tray floor. If the vapor velocity is much lower than design, liquid will begin to flow through the perforations rather than into the downcomer.
This condition is known as weeping. Where weeping is severe, the equilibrium efficiency will be very low. For this reason, sieve trays have a very small turndown ratio.

Figure 6-5. Schematic of a stabilizer tower.

Figure 6-6. Vapor flow through a sieve tray.
Valve Trays
Valve trays are essentially modified sieve trays. Like sieve trays, holes are punched in the tray floor. However, these holes are much larger than those in sieve trays. Each of these holes is fitted with a device called a “valve.” Vapor flowing up through the tower contacts the liquid by passing through valves in the tray floor (Figure 6-7). Valves can be fixed or moving. Fixed valves are permanently open and operate as deflector plates for the vapor coming up through the tray floor. For moving valves, vapor passing through the tray floor lifts the valves and contacts the liquid. Moving valves come in a variety of designs, depending on the manufacturer and the application. At low vapor rates, valves will close, helping to keep liquid from falling through the holes in the deck.
At sufficiently low vapor rates, a valve tray will begin to weep. That is, some liquid will leak through the valves rather than flowing to the tray down-comers. At very low vapor rates, it is possible that all the liquid will fall through the valves and no liquid will reach the down-comers.
This severe weeping is known as “dumping.” At this point, the efficiency of the tray is nearly zero.

Figure 6-7. Vapor flow through valve tray
Bubble Cap Trays
In bubble cap trays, vapor flowing up through the tower contacts the liquid by passing through bubble caps (Figure 6-8).
Each bubble cap assembly consists of a riser and a cap. The vapor rising through the tower passes up through the riser in the tray floor and then is turned downward to bubble into the liquid surrounding the cap. Because of their design, bubble cap trays cannot weep. However, bubble cap trays are also more expensive and have a lower vapor capacity/higher pressure drop than valve trays or sieve trays.

Figure 6-8. Vapor flow through bubble cap tray

High Capacity/High Efficiency Trays
High capacity/high efficiency trays have valves or sieve holes or both. They typically achieve higher efficiencies and capacities by taking advantage of the active area under the down-comer. At this time, each of the major vendors have their own version of these trays, and the designs are proprietary.

Bubble Cap Trays vs. Valve Trays
At low vapor rates, valve trays will weep. Bubble cap trays cannot weep (unless they are damaged). For this reason, it is generally assumed that bubble cap trays have nearly an infinite turndown ratio. This is true in absorption processes (e.g., glycol dehydration), in which it is more important to contact the vapor with liquid than the liquid with vapor. However, this is not true of distillation processes (e.g., stabilization), in which it is more important to contact the liquid with the vapor. As vapor rates decrease, the tray activity also decreases. There eventually comes a point at which some of the active devices (valves or bubble caps) become inactive. Liquid passing these inactive devices gets very little contact with vapor. At this point, it is possible that liquid may flow across the entire active area without ever contacting a significant amount of vapor. This will result in very low efficiencies for a distillation process.
Nothing can be done with a bubble cap tray to compensate for this.
However, a valve tray can be designed with heavy valves and light valves. At high vapor rates, all the valves will be open. As the vapor rate decreases, the valves will begin to close. With light and heavy valves on the tray, the heavy valves will close first, and some or all of the light valves will remain open. If the light valves are properly distributed over the active area, even though the tray activity is diminished at low vapor rates, what activity remains will be distributed across the tray. All liquid flowing across the tray will contact some vapor, and mass transfer will continue. Of course, even with weighted valves, if the vapor rate is reduced enough, the tray will weep and eventually become inoperable.
However, with a properly designed valve tray this point may be reached after the loss in efficiency of a comparable bubble cap tray. So, in distillation applications, valve trays can have a greater vapor turndown ratio than bubble cap trays.

Tray Efficiency and Stabilizer Height
In general, stabilizer trays generally have a 70% equilibrium stage efficiency. That is, 1.4 actual trays are required to provide one theoretical stage. The spacing between trays is a function of the spray height and the down-comer backup (the height of clear liquid established in the down-comer). The tray spacing will typically range from 20 to 30 inches (with 24 inches being the most common), depending on the specific design and the internal vapor and liquid traffic. The tray spacing may increase at higher operating pressures (greater than 165 psia) because of the difficulty in disengaging vapor from liquid in the active areas of the tray.

Packing
Packing typically comes in two types: random and structured. Liquid distribution in a packed bed is a function of the internal vapor/liquid traffic, the type of packing employed, and the quality of the liquid distributors mounted above the packed bed. Packing material can be plastic, metal, or ceramic. Packing efficiencies can be expressed as height equivalent to a theoretical plate (HETP).

Random Packing
A bed of random packing typically consists of a bed support (typically a gas injection support plate) upon which pieces of packing material are randomly arranged (they are usually poured or dumped onto this support plate). Bed limiters, or hold-downs, are sometimes set above random beds to prevent the pieces of packing from migrating or entraining upward. Random packing comes in a variety of shapes and sizes. For a given shape (design) of packing, small sizes have higher efficiencies and lower capacities than large sizes.
Figure 6-9 shows a variety of random packing designs. An early design is known as a Rasching ring. Rasching rings are short sections of tubing and are low-capacity, low-efficiency, high-pressure drop devices. Today’s industry standard is the slotted metal (Pall) ring. A packed bed made of 1-inch slotted metal rings will have a higher mass transfer efficiency and a higher capacity than will a bed of 1-inch Rasching rings. The HETP for a 2-inch slotted metal ring in a stabilizer is about 36 inches. This is slightly more than a typical tray design, which would require 34 inches (1.4 trays × 24-inch tray spacing) for one theoretical plate or stage.

Structured Packing
A bed of structured packing consists of a bed support upon which elements of structured packing are placed. Beds of structured packing typically have lower pressure drops than beds of random packing of comparable mass transfer efficiency. Structured packing elements are composed of grids (metal or plastic) or woven (metal or plastic) or of thin vertical crimped sheets (metal, plastic, or ceramic) stacked parallel to each other. Figure 6-10 shows examples of the vertical crimped sheet style of structured packing.
The grid types of structured packing have very high capacities and very low efficiencies, and are typically used for heat transfer or for vapor scrubbing. The wire mesh and the crimped sheet types of structured packing typically have lower capacities and higher efficiencies than the grid type.

Trays or Packing ?
There is no umbrella answer. The choice is dictated by project scope (new tower or retrofit), current economics, operating pressures, anticipated operating flexibility, and physical properties.

Distillation Service
For distillation services, as in hydrocarbon stabilization, tray design is well understood, and many engineers are more comfortable with trays than with packing. In the past, bubble cap trays were the standard. However, they are not commonly used in this service anymore. Sieve trays are inexpensive but offer a very narrow operating range when compared with valve trays. Although valve trays offer wider operating range than sieve trays, they have moving parts and so may require more maintenance. High capacity/high efficiency trays can be more expensive than standard valve trays. However, high capacity/high efficiency trays require smaller diameter stabilization towers, so they can offer significant savings in the overall cost of the distillation tower. Random packing has traditionally been used in small diameter (<20 inches) towers. This is because it is easier and less expensive to pack these small diameter towers. However, random packed beds are prone to channeling and have poor turndown characteristics when compared with trays. For these reasons, trays were preferred for tower diameters greater than 20 inches.

Stripping Service
For stripping service, as in a glycol or amine contactor, bubble cap trays are the most common. In recent years, there has been a growing movement toward crimped sheet structured packing. Improved vapor and liquid distributor design in conjunction with structured packing can lead to smaller-diameter and shorter stripping towers than can be obtained with trays.

6.3.1.2: Stabilizer Reboiler
The stabilizer reboiler boils the bottom product from the stabilizer tower.
The source of all heat used to generate vapor in a stabilizer is the reboiler. The boiling point of the bottom product is controlled by controlling the heat input of the reboiler together with the stabilizer operating pressure, this actions control the vapor pressure of the bottom product.
Reboiler temperatures typically range from 2000 to 4000F (900 to 2000C) depending on operating pressure, bottom product composition, and vapor pressure requirements. It’s important to note that reboiler temperatures should be kept to a minimum to decrease the heat requirements, limit salt buildup, and prevent corrosion problems.
Maintaining stabilizer operating pressures below 200 psig will result in reboiler temperatures below 3000F. A water-glycol heating medium can then be used to provide heat. Higher stabilizer pressures require the use of steam or hydrocarbon-based heating mediums.
However, operating at high pressures decreases the flashing of the feed when entering the stabilizer tower and decreases the amount of feed cooling required. In general, a liquid hydrocarbon stabilizer should be designed to operate between 100 to 200 psig.
Selection of a stabilizer heat source depends on the medium and tower operating pressure. The source of reboiler heat should be considered when a crude stabilizer is being evaluated. If turbine generators or compressors are installed nearby, then waste heat recovery should be considered.

Figure 6-9. Various types of random packing.

Figure 6-10. Structured packing can offer better mass transfer than trays.

6.3.1.3: Stabilizer Cooler
The stabilizer cooler is used to cool the bottom product leaving the tower before it goes to a tank or pipeline. The temperature of the bottom product may be dictated by contract specification or by efforts to prevent loss of vapors from an atmospheric storage tank.
For a stabilizer with a reflux system, the bottom product may be cooled by exchanging heat with the feed to the stabilizer.

6.3.1.4: Stabilizer Reflux System
The stabilizer reflux system consists of a reflux condenser, reflux accumulator, and reflux pumps. The system is designed to operate at a temperature required to condense a portion of the vapors leaving the top of the stabilizer.
The temperature range is determined by calculating the overhead vapor’s dew point temperature. The heat duty required is determined by the amount of reflux required.
The type of exchanger selected for the reflux depends on the design temperature required to condense the reflux. The lower the operating pressure of the stabilizer, the lower the temperature required for condensing the reflux. In most installations, air-cooled exchangers may be used. Some installations may require refrigeration and a shell-and-tube exchanger configuration.
The reflux accumulator consists of a two-phase separator with several minutes of retention time to allow separation of the vapors and liquids.
The reflux accumulator is normally located below the reflux condenser, with the line sloped from the condenser to the accumulator. The size of the reflux accumulator depends on the amount of reflux required and the total amount of vapors leaving the stabilization tower.
Reflux pumps are sized to pump the required reflux from the reflux accumulator back to the top of the stabilizer tower. Depending upon the reflux circulation rate, two 100 percent pumps or three 50 percent pumps may be installed. This allows either a 100 percent spare or a 50 percent spare pump.

6.3.1.5: Stabilizer Feed Cooler
An inlet feed cooler may be required if a cold feed stabilizer tower is used. Calculations are required to determine the design feed temperature and the heat duty exchanger. This exchanger is usually a shell-and-tube type with some type of refrigerant required to cool the feed sufficiently.

6.3.1.6: Stabilizer-Heater
A feed heater may be required for stabilizers with a reflux system. If a feed heater is used, it is normally a shell-and-tube type exchanger that exchanges heat between the cold feed and the hot bottom product, which is then cooled before going to storage or pipeline.
The selection of equipment and the decision whether to use cold-feed or a reflux system depends on a number of factors. The availability of heat sources for reboiler and streams for cooling the system influence the final decision. Economics of product recovery, CAPEX, and OPEX are major considerations.
6.4: Stabilizer Design
It can be seen from the previous description that the design of both a cold-feed stabilizer and a stabilizer with a reflux is a rather complex and involved procedure. Distillation computer simulations are available that can be used to optimize the design of any stabilizer if the properties of the feed stream and desired vapor pressure of the bottom product are known. Cases should be run of both a cold-feed stabilizer and one with reflux before a selection is made. Because of the large number of calculations required, it is not advisable to use hand calculation techniques to design a distillation process. There is too much opportunity for computational error.
Normally, the contract specification will specify a maximum Reid Vapor Pressure (RVP). This pressure is measured according to a specific American Society of Testing Materials (ASTM) testing procedure.
A sample is placed in an evacuated container such that the ratio of the vapor volume to the liquid volume is 4 to 1. The sample is then immersed in a 1000F liquid bath. The absolute pressure then measured is the RVP of the mixture.
The vapor pressures of various hydrocarbon components at 1000F are given in Table 6-1.
The vapor pressure of a mixture is given by:
VP = ∑ [ VPn x MFn ] Eq. 6-1
Where
Where VP = vapor pressure of mixture, psia
VPn = vapor pressure of component n, psia
MFn = mole fraction of component n in liquid

Table 6-1, Vapor pressure of relative light components.
6.5: Crude Oil Sweetening
Apart from stabilization problems of ‘‘sweet’’ crude oil, ‘‘sour’’ crude oils containing hydrogen sulfide, mercaptans, and other sulfur compounds present unusual processing problems in oil field production facilities. The presence of hydrogen sulfide and other sulfur compounds in the well stream impose many constraints. Most important are the following:
* Personnel safety and corrosion considerations require that H2S concentration be lowered to a safe level.
* Brass and copper materials are particularly reactive with sulfur compounds; their use should be prohibited.
* Sulfide stress cracking problems occur in steel structures.
* Mercaptans compounds have an objectionable odor.
Along with stabilization, crude oil sweetening brings in what is called a ‘‘dual operation,’’ which permits easier and safe downstream handling and improves and upgrades the crude marketability.
Three general schemes are used to sweeten crude oil at the production facilities:

6.6.1: Stage vaporization with stripping gas.
This process—as its name implies—utilizes stage separation along with a stripping agent.
Hydrogen sulfide is normally the major sour component having a vapor pressure greater than propane but less than ethane.
Normal stage separation will, therefore, liberate ethane and propane from the stock tank liquid along with hydrogen sulfide. Stripping efficiency of the system can be improved by mixing a lean (sweet) stripping gas along with the separator liquid between each separation stage. Figure 6-11, represents typical stage vaporization with stripping gas for crude oil sweetening/stabilization. The effectiveness of this process depends on the pressure available at the first-stage separator (as a driving force), well stream composition, and the final specifications set for the sweet oil.

Figure 6-11. Crude sweetening by stage vaporization with stripping gas.

6.6.2: Trayed stabilization with stripping gas.
In this process, a tray stabilizer (nonreflux) with sweet gas as a stripping agent is used as shown in Figure 6-12. Oil leaving a primary separator is fed to the top tray of the column countercurrent to the stripping sweet gas. The tower bottom is flashed in a low-pressure stripper. Sweetened crude is sent to stock tanks, whereas vapors collected from the top of the gas separator and the tank are normally incinerated. These vapors cannot be vented to the atmosphere because of safety considerations. Hydrogen sulfide is hazardous and slightly heavier than air; it can collect in sumps or terrain depressions.
This process is more efficient than the previous one.

Figure 6-12. Crude sweetening by trayed stabilization with stripping gas.
6.6.3: Reboiled trayed stabilization.
The reboiled trayed stabilizer is the most effective means to sweeten sour crude oils. A typical reboiled trayed stabilizer is shown in Figure 6-13. Its operation is similar to a stabilizer with stripping gas, except that a reboiler generates the stripping vapors flowing up the column rather than using a stripping gas. These vapors are more effective because they possess energy momentum due to elevated temperature.
Because hydrogen sulfide has a vapor pressure higher than propane, it is relatively easy to drive hydrogen sulfide from the oil.

Figure 6-13. Crude sweetening by reboiled trayed stabilization.

yasserkassem · Post by **yasserkassem** » Sun Mar 22, 2020 1:34 pm

Chapter 7 -
Oil and Gas Measurements

-------
Chapter 7 209
Fluid Measurements 209
7.1: Gas Measurement 209
7.1.1: Orifice-Meter Measurement 209
7.1.1.5: Meter Tubes 213
7.1.2: Ultrasonic Measurement 220
7.2: Liquid Measurements 221
7.2.1: Volumetric Measurement Meters (Orifice Meters) 221
7.2.2: Turbine Meters 223
7.2.3: Positive Displacement Meters 224
7.2.4: Turbine and Positive Displacement Meter Selection 224
7.2.5: Mass Measurement Meters 225
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Chapter 7

Fluid Measurements
7.1: Gas Measurement
7.1.1: Orifice-Meter Measurement
The most commonly used differential measurement device, the orifice meter, is widely accepted for use in measuring volumes of liquids or vapors.
The orifice meter consists of static pressure and differential pressure recording gauges connected to an orifice flange or orifice fitting. The orifice meter tube (meter run) consists of upstream and downstream sections of pipe for which size and tolerance have been determined through calculation and which conform to specifications set forth in API Chapter 14.3.
The orifice plate is held perpendicular to flow by flanges or a fitting. Bore, circumference, edge sharpness, and other tolerances must meet specifications as set forth in API Chapter 14.3.
For additional calculation procedures, refer to in Chapter 14.3 of the API Manual of Petroleum Measurement Standards.

7.1.1.1: Orifice Flanges (Fig. 7-1a)
Orifice flanges require that the line be shut down and depressurized in order to inspect or change the orifice plate. The flange bolts are loosened and removed.
The flanges are spread by use of "jack" bolts, and the plate is removed.
When orifice flanges are used, the pressure tap hole placement may be determined by measuring from the face of the flange to the pressure tap hole center.

7.1.1.2: Single Chamber Orifice Fitting (Fig. 7-1b)
This fitting also requires that the line be shut down and depressurized in order to inspect or change the orifice plate. However, this fitting does not require breaking apart the flanges. Instead, the bolts are loosened on the cover plate and the cover plate removed. The orifice plate holder and orifice plate are then removed from the fitting. These fittings provide precise alignment of the orifice plate.

7.1.1.3: Dual-Chamber Orifice Fitting (senior fitting) (Fig. 7-1c).
This fitting allows the removal and inspection of an orifice plate while the line remains under pressure. It allows the orifice plate holder and orifice plate to be raised into the upper cavity of the fitting by the use of a crank handle. A valve is then closed to separate the upper cavity from the lower cavity of the fitting. The upper cavity is then depressurized, the top cover plate removed, and the orifice plate cranked out.

7.1.1.4: Orifice Plates
The minimum, maximum, and recommended thicknesses of orifice plates for various pipe sizes are given in Table 7-1. Also shown in this figure are maximum allowable differential pressures for stainless steel plates of the recommended thickness at a maximum temperature of 150 degrees F.
The thickness of the orifice plate at the orifice edge (e) shall meet the following:
The minimum thickness is defined by e ≥ 0.01d or e > 0.005 in. whichever is greater.
The maximum thickness is defined by e ≥ 0.02d or ≤ 0.125d whichever is smaller
If the thickness of the orifice plate must be greater than permitted by these limitations, the downstream edge shall be cut away (beveled or recessed) at an angle of 45°± 15° or less to the face of the plate, leaving the thickness of the orifice edge within these requirements. All orifice plates which are beveled should have the square-edge side (i.e., the side opposite the beveling) stamped "inlet" or the beveled side stamped "outlet".
The upstream face of the orifice plate shall be flat and perpendicular to the axis of the meter tube, when in position between the orifice flanges or in the orifice fitting. Any plate that does not depart from flatness along any diameter by more than 0.010 inch per inch of the dam height, (D-d)/2, shall be considered flat. See Figure 7-2, table. 7-1.
The upstream edge of the orifice shall be square and sharp so that it will not show a beam of light when checked with an orifice edge gauge, or alternately will not reflect a beam of light when viewed without magnification. The orifice shall not have a burred or feathered edge. It shall be maintained in this condition at all times. Moreover, the orifice plate shall be kept clean at all times and free from accumulation of dirt, ice, and other extraneous material. Orifice plates with small nicks in the edge can be expected to increase the flow measurement uncertainty.

Fig. 7-1. Orfice flange and orifice fitting.

Fig. 7-1. Continued.

Fig. 7-2. Orifice Plate Dimensions

Table. 7-1. Orifice Plate Dimensions
Notes:
1. The maximum allowable ΔP (In. of H2O) orifice flange is 1000 for all orifice sizes.

7.1.1.5: Meter Tubes
The term "meter tube" shall mean the straight upstream pipe ahead of the orifice fitting of the same internal diameter as the orifice fitting (length UL on Fig. 7-3) and the similar downstream pipe (length DL on Fig. 7-3) following the orifice.
The sections of pipe to which the orifice flanges are attached or the sections adjacent to the orifice flange or fitting shall comply with API Chapter 14.3 (AGA Report No. 3).
See Tables 7-2, 7-3, and 7-4 for proper meter tube lengths.

FIG. 7-3. Orifice Meter Tube Layout for Flanged or Welded Limit

Table. 7-2.Orifice Meter Installation Requirements without a flow conditioner. Minimum straight unobstructed meter tube length from the upstream and downstream side of the orifice plate (in internal pipe diameter, Di)

Table. 7-2. (continued) Orifice Meter Installation Requirements without a flow conditioner.
UL – Minimum meter tube length upstream of the orifice plate in internal pipe diameter, Di (See Fig. 7-3). Straight length shall be measured from the downstream end of the curved portion of the nearest (or only) elbow or of the tee or the downstream end of the conical portion of reducer or expander
DL – Minimum downstream meter tube length in internal pipe diameters, Di (See Fig. 7-3).
Sep – Separation distance between piping elements in internal pipe diameter, Di, measured from the downstream end of the curved portion of the upstream elbow to the upstream end of the curved portion of the downstream elbow.
Note: The tolerance on specific length for UL and DL is +/- 0.25Di.

Table. 7-3. Orifice Meter Installation Requirements with 1998 uniform Concentric 19-Tube Bundle Flow Straightener for Meter Tube Upstream Length of 17Di =< UL<29Di.

Notes:
Length shown under the UL2 column are the dimension shown in figure 7-3 expressed as the number of published internal pipe diameters (Di) between the downstream end of the 1998 Uniform Concentric 19-Tube Bundle Flow Straightener and the upstream surface of the orifice plate.
( * ) – 13 Di allowed for up to β = 0.54
( ** ) – 9.5 Di allowed for up to β = 0.47
( *** ) – 9.5 Di allowed for up to β = 0.46
Sep – separation distance defined in previous table.
UL1 = UL – UL2
Note: The tolerance on specific length for UL, UL2, and DL is +/- 0.25Di.
Not allowed means that it is not possible to find an acceptable location for the 1998 Uniform Concentric 19-Tube Bundle Flow Straightener downstream of the particular fitting for all values of UL.

Table. 7-4. Orifice Meter Installation Requirements with 1998 Uniform Concentric 19-Tube Bundle Flow Straightener for Meter Tube Upstream Length of UL=>29Di.
Notes:
Length shown under the UL2 column are the dimension shown in figure 7-3 expressed as the number of published internal pipe diameters (Di) between the downstream end of the 1998 Uniform Concentric 19-Tube Bundle Flow Straightener and the upstream surface of the orifice plate.
Sep – separation distance defined in previous table.
UL1 = UL – UL2
Note: The tolerance on specific length for UL, UL2, and DL is +/- 0.25Di.
Not allowed means that it is not possible to find an acceptable location for the 1998 Uniform Concentric 19-Tube Bundle Flow Straightener downstream of the particular fitting for all values of UL.
7.1.1.6: Flow Conditioners (Fig. 7-4)
The purpose of flow conditioners is to eliminate swirls and cross currents set up by the pipe, fittings and valves upstream of the meter tube.
Please refer to API Chapter 14.3 (AGA Report No. 3) for detailed specifications for flow conditioners.

FIG. 7-4. 1998 Uniform Concentric 19-Tube Bundle Flow Straightener

7.1.1.7: Gas Orifice Calculations
Orifice Sizing
A simple calculation is often needed to properly size an orifice plate for new or changing flow rates through existing meter tubes. The procedure uses an existing or assumed flow quantity, a desired differential pressure at a specific static pressure, an estimated flowing temperature, and a determined or assumed specific gravity. The Key orifice coefficient is calculated from the gas flow equation. This calculated value is then compared to Table. 7-5, and the next larger size is usually selected.

To determine the approximate orifice size required, the corresponding Keyg (natural gas) is calculated using appropriate terms of Eq. 7-1;
Qh = Keyg x Ftf x Fg (hw x Pf)0.5 Eq. 7-1

Where
Qh = rate of flow, std. cu ft/hr
Keyg = Fn (Fc + Fsl) = orifice factor
Ftf = flowing temperature factor to change from the assumed flowing temperature of 60°F to the actual flowing temperature = [520/(460+Tf)]0.5
Fg = specific gravity factor applied to change from a specific gravity of 1.0 (air) to the specific gravity of the flowing gas = (1/ G gas Sp. Gr.)0.5
hw = differential pressure measured across the orifice plate in inches of water at 60°F (1 psi = 2.77 inches of water)
Pf = flowing pressure psia.

Table. 7-5. Plate Sizing and Approximate Flow rate, Natural Gas, Natural Gas Liquids and Steam.
(For Flange Taps).

Example 7-1: Size an orifice plate in gas service.
Given Data:
Line Size, D = 4.026 in.
Flange Taps
Specific Gravity = 0.700
Flowing Temperature = 100°F
Flowing Pressure = 75 psia
Flow Rate = 14,200 cu ft/hr
(14.73 psia @ 600F)
Desired Differential = 50 in. of water

Solution (using equation 7-1)
Ftf = (520/560)0.5 = 0.9636
Fg = (1/0.7)0.5 = 1.1952

14,200 = Keyg x 0.9636 (1.1952) (50 x 75)0.5
Keyg = 201.342

Referring to Keyg (Table. 7-5) for a 4.026 inch line with flange taps, access the Keyg value which approximates the calculated number. A 1.000 in. orifice size would be selected which has a
Key value of 201. More precise calculations would include other corrections. For more precise custody transfer calculations, please refer to API Chapter 14.3 (AGA Report No. 3).

Flow Rate Calculation
The following example illustrates a calculation of flow rate through an orifice.

Example 3-2: Calculate an approximate flow rate for the orifice using appropriate terms from Eq 7-1.
Given Data:
Line Size, D = 6.065 in.
Orifice Size, d = 3.500 in.
Flange Taps
Flowing Temperature = 70°F
Flowing Pressure = 90 psia
Differential = 60 in. of water
Specific Gravity = 0.750

Qh = Keyg x Ftf x Fg (hw x Pf)0.5 Eq. 7-1
From table 7-5, Keyg = 2646
Ftf = (520/530)0.5 = 0.9905
Fg = (1/0.75)0.5 = 1.1547

Qh = 2646 (0.9905) (1.1547) (60 x 90)0.5

Qh = 222,387 cu ft/hr @ 14.73 psia and 60°F

More precise calculations would include other corrections.
For more precise custody transfer calculations, please refer to API Chapter 14.3 (AGA Report No. 3).
Well Test Calculation
Often it is necessary to determine an approximate flow quantity from a well head or field separator vent to the atmosphere for test purposes. The use of a "well head tester" has been a common practice since the early days of the oil and gas industry. See Figure 7-5. An orifice is installed between a pair of flanges, at the outlet of a pipe nipple which is at least eight pipe diameters long. The square edge of the orifice faces the flow. The diameter of the pipe nipple should not be greater than the preceding fittings. The pressure connection may be made in the upstream flange or at any point in the pipe nipple within three diameters from the orifice. The pressure differential across the orifice is the difference between the upstream pressure and atmospheric pressure.

FIG. 7-5. Typical Test Set-Up for Measuring Gas from a Separator Vent

An approximate flow rate may be calculated from:

Q = 16,330 (1 + β4) (d2) Ftf x Cg [H(29.32+0.3H)]0.5 Eq. 7-2

For conditions other than 60°F (flowing) and G of 0.6, correction factors must be applied.
Ftf = [520/(460+Tr)]0.5 Eq. 7-3
Cg = (0.6/ G gas Sp. Gr.)0.5 Eq. 7-4

Where
Q = gas flow rate, scfd
β = ratio of the orifice or throat diameter to the internal diameter of the meter run, dimensionless
d = orifice diameter, in.
Ftf = flowing temperature factor to change from the assumed flowing temperature of 60°F to the actual flowing temperature
Cg= gravity correction factor for orifice well tester to change from a gas specific gravity of 0.6
H = pressure, inches of mercury (1 psi = 2.04 inches of mercury)
Tf = flowing temperature, °F
G = specific gravity at 60°F

Example 3-3 — Calculate the daily gas flow through a 1-inch orifice in a nominal 3-inch pipe. The gas gravity is 0.70, the flowing temperature is 60°F, and the pressure upstream of the orifice is 5 inches Hg. The published ID of a 3-inch pipe is 3.068 in.
Solution:
β = 1/3.068 = 0.3259
β4 = 0.01128
1 + β4 = 1.01128
Ftf = 1
Cg= (0.6/0.7)0.5 = 0.9258

Q = 16,330 (1 + β4) (d2) Ftf x Cg [H(29.32+0.3H)]0.5 Eq. 7-2

Q = 16,330 (1.01128) (1) x 1 x 0.9258 [5(29.32+1.5)]0.5

= 190,000 scfd.

7.1.2: Ultrasonic Measurement
This section gives a short overview of ultrasonic meters. If meter design and custody quality calculations are required, please refer to American Gas Association Report #9. These meters are designed for measurement of single-phase fluid only.
An ultrasonic meter (UM) is a fluid velocity-sensing device.

Fig. 7- 6. Ultrasonic Flow Meter

(See Figure 7-6) The flowing gas velocity is determined by the transit times of high frequency pulses between two matched transducers. One is designated as upstream and one as downstream to the position in the meter and the direction of flow. These transducers attach into the pipe wall but do not protrude into the gas stream, thus creating a zero pressure drop. There are simple, single path meters that consist of one pair of transducers and multi-path meters with three or more pairs of transducers. Each pair of transducers measures the transit time of each sound pulse transmitted from the up-stream transducer to the downstream transducer with the flow (t1), and from the downstream to the upstream transducers against the flow (t2). The transit time for a signal traveling with the gas flow is less than travel time against the gas flow. The difference in these transit times relates to the gas velocity along that specific path. Various calculations and methodologies are then used to calculate the average gas velocity and flow rate at line conditions.
A single path meter monitors only one path’s mean velocity at one elevation in the gas flow. Since most gas flow is not fully symmetrical, the use of a single path UM would have inaccuracies dependent on the flow velocity profile. Single path meters are generally used for operational balancing and flare measurement and are generally not accepted for custody transfer measurement.
A multi-path UM continuously monitors three or more mean velocities at different elevations in the gas stream of the metered area. The averages of these mean velocities are used to calculate the gas flow rate. Meter designs of various meter manufactures are able to minimize the effect of non-symmetrical flow profiles on the overall meter accuracy. It is recommended that UM’s with three or more paths be used for custody measurement (based on available data).

7.2: Liquid Measurements
7.2.1: Volumetric Measurement Meters (Orifice Meters)
Liquid volume measurement by an orifice meter can be determined by following the guidelines established in API Chapter 14.8. As with gas measurement, the primary element should consist of an orifice plate, the orifice holder with it’s associated tap holes to sense the differential and static pressure, and the upstream and downstream piping “meter tube”.
The differential and static pressure readings are sensed at the flange taps by a secondary element sensor or transducer. The temperature of the fluid should also be recorded by the temperature sensor or transducer. Note that the meter is the tertiary device that records the output of the sensors/transducers.
The Reader-Harris, Gallagher equation used with orifice meters produces discharge coefficients accurate within +/- 0.5%. Measurement using orifice meters must include this uncertainty, as well as the uncertainty in the metering equipment, unless the metering system is proven against a traceable standard (see API Chapter 4), similar to the way turbine meters and PD meters are typically proven. Then the overall system uncertainty may be reduced to +/- 0.25%.
Some fluid physical properties also need to be known. Examples may include density, viscosity, and compressibility to accurately determine volume using the AGA Report #3 method. For systems performing custody transfer mass measurement for light hydrocarbons such as ethane, ethylene, E/P Mix, high ethane raw mix NGLs, etc., the flowing density of the stream should be measured with a density meter. Then the mass of the delivery may be determined by multiplying the volume at flowing conditions from the meter/ELM, times the density of the flowing stream from the density meter. Details of this method can be found in API Chapters 14.4, 14.6, 14.7, 14.8, and 21.2.

The following equation is to determine orifice size required, or liquid flow rate,.
Qh = Keyl x Fgt (hw)0.5 Eq. 7-5
Where
Qh = rate of flow, gal./hr
Keyl = Fn (Fc + Fsl) = orifice factor
Fgt = gravity-temperature factor for liquids

Fgt = [1.0057/ (Gl)0.5]x (Gf/Gl)0.5 Eq. 7-6

hw = differential pressure measured across the orifice plate in inches of water at 60°F
Gf = specific gravity at flowing temperature (Extracted from tables of specific gravity correction factors if flowing temperature differ than 600F or from fig 7-7, Extract API value at flowing temperature, and convert it to specific gravity)
Gl = specific gravity at 60°F

Example 3-4: Calculate an approximate orifice size for the given flow rate and line size.
Line Size, D = 3.068 in.
Flange Taps
Specific Gravity at 60°F = 0.690
Flowing Temperature = 60°F
Flow Rate = 3400 gal/hr.
Desired Differential = 50 in. of water

Solution:
Fgt = [1.0057/ (Gl)0.5]x (Gf/Gl)0.5
Fgt = [1.0057/ (0.69)0.5]x 1
= 1.2107
Qh = Keyl x Fgt (hw)0.5
3400 = Keyl x 1.2107(50)0.5
Keyl = 397

Referring to the Key values (Table. 7-5) for a 3.068 inch line with flange taps, access the value listed which approximates the calculated Keyl. A 1.5 inch orifice diameter would be selected, which has a 471 Keyl value.

Flow Rate calculation
The liquid flow rate through an orifice is calculated using eq. 7-5.
The initial calculation can be completed using only the Keyl and the Fgt correction factors to solve for Qh since those factors are most significant.

Example 3-5 — Calculate a liquid flow rate for the given orifice setting.
Line Size, D = 8.071 in.
Orifice Size, d = 4.000 in.
Flange Taps
Specific Gravity at 60°F = 0.630
Flowing Temperature = 80°F
Differential = 36 in. of water

Solution:
Qh = Keyl x Fgt (hw)0.5 Eq. 7-5

The value of Keyl from table. 7-5 is 3345 for an 8.071 in. line with a 4.0 in. orifice. The value of Fgt is calculated using eq. 7-6.
Fgt = [1.0057/ (Gl)0.5]x ( Gf/Gl)0.5 Eq. 7-6

API =89 at 80 0F from fig. 7-7
Sp.gr at 80 0F (Gf) = 141.5/220.5 = 0.6417
Fgt = [1.0057/ (0.63)0.5]x ( 0.6417/0.63)0.5
Fgt =1.267 x 1.009 = 1.278
Therefore,
Qh = 3345 x 1.278 (36)0.5 = 25,649 gal/hr.
For more precise calculations, refer to Chapter 14.8 of the API Manual of Petroleum Measurement Standards.

Figure 7-7. Specific gravity of petroleum fractions. kW is Watson characterization factor, (use approximate value 11, if data is not available), for more precious calculations use API correction tables for temperature.

7.2.2: Turbine Meters
Turbine meters are velocity-sensing devices. The direction of flow through the meter is parallel to a turbine’s rotary axis and the speed of rotation of the rotor is proportional to the rate of flow.
The turbine meter normally consists of one moving part; an impeller held in place by high pressure, low drag bearings. A magnetic transducer mounted in the meter body is used to count revolutions as the flow passes. The pulses from the transducer are determined for a known volume passing through the meter to develop a factor in pulses per gallon, or other desired unit volume. Turbine meter components are shown in Fig. 7-8. Expected accuracies of plus or minus 0.25% can be attained by certain turbine meters where proper stream conditions are maintained and the meter is properly installed and proven.
Doing mass measurement with turbine meters is often preferred where conditions in temperature, pressure, intermolecular adhesion and solution mixing present difficulty in converting volumes from flowing conditions to standard conditions, such as with ethane, natural gas liquids (NGL), or ethane-propane mixes. To do this properly an online densitometer needs to be used. Refer to GPA 8182 or API Chapter 14, Section 7 (14.7) for further details on mass measurement for NGLs.

7.2.3: Positive Displacement Meters
Displacement meters take a physically enclosed volume of fluid and move it from upstream to downstream of the metering point. The sum of these operations is an indication of the amount of liquid, which is moved over a period of time.
An expected accuracy of 0.25% for a positive displacement (PD) meter can be attained when it is properly installed and proven. Application is normally limited to those fluids that exhibit some lubricating properties because of the multiple moving parts of a positive displacement meter. Typical applications are butane and heavier products since ethane and propane have minimal lubricating properties. Fig. 7-9 shows some internal details of a positive displacement meter. PD meters may perform mass or volumetric measurement, depending on their configuration and companion equipment.

Fig. 7-8. Turbine meter.

7.2.4: Turbine and Positive Displacement Meter Selection
Turbine and positive displacement meter installations should include the following considerations:
• Application to proper flow ranges
• Upstream strainers to protect meter internals from foreign material
• Pulsation and vibration
• Proper upstream flow conditioning
• Significant rate changes
• Changes in flow temperature, pressure, and density
• Back pressure (2 times “ DP” across meter plus 1.25 times equilibrium vapor pressure is minimum recommended). [“DP” is the difference between the flowing pressure and the equilibrium vapor pressure of the liquid.]
• Connections to prove the meter
• Verification that Liquid temperature correction factor (Ctl) and liquid pressure correction factor (Cpl) will not be applied when the meters are performing mass measurement, except during provings.
The normally acceptable performance of a turbine or positive displacement meter will result in a change in the pulse count of less than 0.05% between meter prover runs, and less than 0.25% between provings. If the factor changes more than 0.25% between provings:
• meter maintenance may be required
• a total flow adjustment must be made
If the factor changes more than 0.5% between provings:
• the turbine must be pulled and inspected for damage or wear
• a total flow adjustment must be made
• the turbine must be proven again following inspection
More details about turbine and positive displacement meter installations, operation, and proving are available in Chapters 4, 5, 6, and 12 of the API Manual of Petroleum Measurement Standards.

7.2.5: Mass Measurement Meters
Mass measurement of a flowing fluid is advantageous where the physical properties of the fluid are not well defined or available. Mass measurement is especially important in measuring streams containing ethane and methane because of substantial solution mixing effects. Mass measurement is accomplished by multiplying the volume of the fluid at flowing conditions, over a defined period of time, by the density of the fluid at flowing conditions during that same time. This procedure eliminates the need for the correction factors (Ctl and Cpl) for the metered volume. The total stream mass can be converted into pure components by using a weight analysis of the fluid. Refer to GPA 8182 or API Chapter 14, Section 7 (14.7) for further details on mass measurement for NGLs.
Several more different techniques and processes have been developed to directly measure the mass of a flowing fluid. The devices utilize the principle that angular momentum of a mass is directly proportional to the mass velocity. The resistance of a mass to change direction is measured by different types of devices using combinations of sensitive mechanical and electrical sensors and transmitters that can result in a variety of electronic signals. Mass flow meter installations may not require upstream and downstream piping usually associated with other types of measurement. Proving mass flow meters may involve a complicated arrangement of flow and density measuring equipment, or access to an alternate proving station, or use of a master mass meter comparison.

7.2.5.1: Coriolis Meters
there is no flow, the two sine waves produced are in phase. When there is flow, the induced Coriolis force causes the tubes to twist, resulting in two out-of-phase sine waves.
The time difference in the sine waves is directly proportional to the mass flow rate through the tubes (this The Coriolis meter is a mass-measuring device. It consists of a sensor, a transmitter and peripheral devices to provide monitoring, alarm, and/or control functions.
The sensor consists of two flow tubes, the drive coil and magnet, two pick-off coils and magnets and the RTD “Resistance Temperature Devices.”. During operation, process fluid entering the sensor is split, half passing through each flow tube. The drive coil is energized causing the tubes to oscillate up and down in opposition to one another.
The pick-off coils are mounted on one tube while the magnets are mounted on the other. Each coil moves through the uniform magnetic field of the adjacent magnet as the two tubes move. The voltage generated from each pickoff coil creates a sine wave representing the motion of one tube relative to the other. When may only be true at a fixed pressure).
The density of the fluid is calculated from the frequency of oscillation of the tubes.
The transmitter provides three actions. First, it sends a pulsed current to the sensor drive coil causing the flow tubes to vibrate. Second, it processes the sensor input signals, performs calculations, and produces various outputs to peripheral devices. Most commonly, the output of the meter is a pulsed output. Third, it allows communication with an operator or control system. Figure 7-10 shows the components of a Coriolis meter.

Fig. 7-9. Positive displacement meter.

FIG. 7-10. Components of a Coriolis Meter
Detail description of how a Coriolis meter operates can be found in appendix A of the API Coriolis Liquid Measurement Draft Standards.
When configuring the meter, users should provide some means to block in the flow so the zero flow condition can be verified. Zero verification of the meter is required from time to time as part of the normal operating procedures. Zeroing is necessary when the zero offset has shifted outside the defined limits. Since the meter should be proven after each zero, unnecessary zeroing should be avoided to minimize potential errors associated with meter factor reproducibility.
The Coriolis meter should be proven under conditions as close to normal operating conditions as practical. The result of a meter proving will be a new or reaffirmed meter factor (MF). This meter factor may be entered in accessory equipment, the Coriolis transmitter, or applied manually to the quantity indicated. The preferred method is to input the meter factor into the accessory equipment due to its audit trail capabilities. A Coriolis meter is normally set up with calibration factors from the manufacturer. These factors, although adjustable, should not be changed. Figure 7-11 shows a typical schematic of a Coriolis meter installation. For more information on Coriolis meter, please refer to the API draft standard, Measurement of Single-phase Intermediate and Finished Hydrocarbon Fluids by Coriolis Meters.

FIG. 7-11. Typical Installation of a Liquid Coriolis Meter

The Coriolis meter has gained popularity in recent years as it presents a number of advantages over other types of meters. A Coriolis meter has an accuracy range of (+/-0.1%) and acceptable repeatability. It provides multi-variable measurement in one device: mass flow rate, volumetric flow rate, density and temperature. It is very tolerant of the changes in the fluid quality and flow rate. It may also be used as a bi-directional meter. Ease of installation and low maintenance are other bonuses as there are no special mounting, no flow conditioning, no straight pipe run requirements and no moving parts.
Like all other types of meters, the Coriolis meter has its own down side. There is a significant pressure drop across the meter making it unsuitable for an existing operation where additional pressure drop cannot be tolerated.

yasserkassem · Post by **yasserkassem** » Sun Mar 22, 2020 1:38 pm

Chapter - 8

Instrumentation and Control

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Chapter 8 228
Instrumentation and Control 228
8.1: Introduction 228
8.2: Type Selection and Identification 228
8.2.1: Pneumatic Power Supplies 228
8.2.2: Electronic Power Supplies 229
8.3: Sensing Devices 230
8.3.1: Pressure Sensors 230
8.3.1.3: Bellows (Fig. 8-3) 230
8.3.2: Level Sensors 232
8.3.3: Temperature Sensors 237
8.3.4: Flow Sensors 239
8.4: Signal Transmitters 241
8.4.1: Pneumatic Transmitters 241
8.4.2: Electronic Transmitters 241
8.5: Signal Converters 241
8.5.1: Pneumatic-to-electronic (P/I) 242
8.5.2: Electronic-to-pneumatic (I/P) 242
8.5.3: Isolators 242
8.5.4: Electric signal converters 242
8.5.5: Frequency converters 242
8.6: Recorders and Indicators 242
8.6.1: Recorders 242
8.6.2: Indicators 242
8.7: Control Concepts 243
8.7.1: Control Loops 243
8.8: Control Modes and Controllers 245
8.8.1: Two-Position (on-off) Controllers 245
8.8.2: Proportional Control Mode 245
8.9: Control Valves 246
8.9.1: Control-Valve Bodies 247
8.9.2: Control-Valve Actuators 248
8.9.3: Flow Characteristics and Valve Selection 249
8.9.4: Fundamentals of Control Valve Sizing 250

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Chapter 8

Instrumentation and Control

8.1: Introduction
Instrumentation in oil and gas processing plant is usually comprised of a system of pneumatic, hydraulic, and electronic devices for measurement and control of all the process variables (pressure, flow, temperature, etc.) which are pertinent to the operation of the plant. In addition, computers are normally included in the instrumentation system to handle functions such as data gathering and transmission, bulk data storage, display, alarms, logging, computations, and control. Since the advent of integrated circuit electronics, specifically the microprocessor, many types of instruments are becoming more intelligent or “computerized.” An instrument may perform a single function such as a temperature indicator (TI), or a combination of functions such as a flow recording controller (FRC).
8.2: Type Selection and Identification
Often the type selection of an instrument is pre-determined by whatever is available, or what will be compatible with the rest of a system. There are cases, however, where the choice to install pneumatic or electronic instrumentation must be made by comparing the features of each type.

8.2.1: Pneumatic Power Supplies
The pneumatic power supply is more commonly known as the instrument air system. The main considerations of an instrument air system are:
1. Adequate Capacity: The minimum capacity of the system should be the sum of the individual requirements of each air-consuming instrument in the system, plus a supplemental volume for purges, leaks, additions, etc. If accurate consumption figures are not available, an estimated consumption volume of 0.5 cubic foot per minute for each air-consuming device is usually adequate.
The air storage tank should have sufficient capacity to maintain this rate for about five minutes or such time as is considered adequate to perform an emergency shut-down of the plant or to switch over to a backup air system. Also the air storage tank capacity should be large enough to prevent excessive cycling of the compressor.
2. Filtering and Regulation: Instrument air systems are normally designed for pressures up to 125 psig and should be protected by relief valves. Instrument air should be free from all contamination such as oil, water, and any hazardous or corrosive gases. Non-lubricated compressors should be used if possible. Where lubricated compressors are used, an oil removal separator is required. The presence of oil may cause instrument contamination and possibly create a combustible mixture. After being compressed, instrument air must be cooled to remove the major portion of the contained water. A final drying system must be used to reduce the water dew point to at least 10°F below the ambient temperature at line pressure. An after filter may be required to remove particulate carryover from the dehydrators.
3. Proper Distribution: The air distribution system should be free of any “pockets” where liquid could accumulate. If this is not possible, drain valves should be installed.
All supply lines should connect to the top of the air manifold or “header.” Instrument air filter-regulators should be provided at each air-consuming device to reduce the line pressure to the supply pressure recommended by the instrument manufacturer. Instrument Society of America Standards ISA-S7.3 and ISA-S7.4 are references for additional information.
4. Non-Air Systems: Natural gas has been used instead of instrument air in some remote installations where compressed air was not available. This practice should be avoided if at all possible due to safety and pollution problems and the additional filtering and clean-up of the gas which must be done to protect the instruments. The user must be cognizant of all applicable regulations when considering the use of any combustible gas in instrumentation service. Some small-scale systems have used bottled nitrogen for instrument gas. This is quite acceptable, but non-bleed type instruments should be used to keep the consumption to a minimum.
5. Hydraulic Powered Devices: Hydraulic actuators are sometimes used on valves or rams where very high thrusts (up to 50,000 pounds force) are required for operation.
Due to the problems of transmitting very high pressure signals, a local pump powered by an electric motor is often used to form what is commonly known as an “electro-hydraulic actuator.”

8.2.2: Electronic Power Supplies
Electronic systems send and transfer signals in cables distributed all over the process plant, working in a voltage range 0-1 Volt and current density 4-20 mA. Specifications and codes related to electronic system is more specific for electrical and instrumentations engineers and cannot be presented here.
In the following, table a comparison between pneumatic and electrical process control.
Pneumatic Electronic
Advantages
1. Intrinsically safe, no electrical circuits. 1. Greater accuracy.
2. Compatible with valves. 2. More compatible with computers.
3. Reliable during power outage for short period of time, dependent on size of air surge vessel. 3. Fast signal transit time.

4. No signal integrity loss if current loop is used and signal is segregated from A.C. current.
Disadvantages
1. Subject to air system contaminants. 1. Contacts subject to corrosion.
2. Subject to air leaks.
2. Must be air purged, explosion proof, or intrinsically safe to be used in hazardous areas.
3. Mechanical parts may fail due
to dirt, sand, water, etc. 3. Subject to electrical interference

4. Signal boosters often needed on transmission lines of over 300 feet. 4. More difficult to provide for positive fail-safe operation.
5. Subject to freezing with moisture present. 5. Requires consideration of installation details to minimize points 1, 2, 3, and 4.
6. Control speed is limited to velocity of sound.

Table. 8-1. Instrument Type Features.

8.3: Sensing Devices
Some of the more common types of sensing devices for the measurement of process variables are described as follows:
8.3.1: Pressure Sensors
8.3.1.1: Manometer (Fig. 8-1)
Two different pressures are applied to two separate openings in a transparent vessel containing a liquid. The difference in the heights of the liquid is used as a measure of the differential pressure. This difference should be corrected for temperature and gravity of the liquid in the manometer (usually either water or mercury). Pressures are often expressed in units such as “inches of water” or “millimeters of mercury.”

FIG. 8-1. Types of Manometers

8.3.1.2: Bourdon tubes (Fig. 8-2)
A Bourdon tube is a metallic coil constructed from a metal tube having the desired elastic quality and corrosion resistance. The tendency of the tube to straighten under pressure causes a mechanical linkage to move a pointer or initiate pneumatic or electronic transmission of the measured pressure. Dampeners should be used where pulsation is a problem. Condensate traps should be used upstream of the device in steam service. The pressure indicated is “gauge” pressure which is relative to that of the surroundings. Bourdon gauges are also available as “compound” types which indicate vacuum as well as positive pressure.

8.3.1.3: Bellows (Fig. 8-3)
A tubular device with pleated sections somewhat like an accordion. It is flexible along its axis and lengthens or shortens according to the applied pressure.
The bellows is usually used in low pressure or vacuum service but types are available for use with high pressures (up to several thousand psi). Typical diameters range from 1/2" to 12".
They are often used in force-balance type transmitters and other applications where small displacements are required. Like the Bourdon tube, it indicates pressures as “gauge” or relative to its surroundings.

FIG. 8-2. Types of Bourdon Tubes

FIG. 8-3. Types of Bellows

8.3.1.4: Diaphragm (Fig. 8-4)
A flat or curved seal with a link attached to an indicator or transmission device. A diaphragm may have its own deflection properties such as with a metallic type or it may be attached to a spring or other elastic member such as with non-metallic diaphragms.

8.3.1.5: Electrical Pressure Transducers
The primary sensing element of many electrical pressure transducers usually takes the form of a Bourdon tube, bellows, or diaphragm to generate a movement which is transmitted to a strain gauge. A strain gauge is a device using resistance wire connected in a Wheatstone bridge configuration to generate an electrical signal proportional to the movement and hence proportional to the process variable being measured. Other types of electrical pressure transducers use properties of inductance, capacitance, or magnetic coupling to convert a pressure measurement to an electrical signal.

FIG. 8-4. Pressure diaphragm elements.

8.3.2: Level Sensors
8.3.2.1: Gauge glass (Fig. 8-5)
This is the most commonly used visual process-level device. Gauge glasses are generally classified as either transparent or reflex types. A transparent gauge glass consists of either a glass tube or an arrangement of flat glass plates in some type of holder. Since the process fluid level is viewed directly, the transparent gauge glass is normally used with opaque fluids. The reflex type has reflecting prisms to aid in viewing transparent fluids.

FIG. 8-5. Flat glass gauge glasses

8.3.2.2: Chain and tape float gauges (Fig. 8-6)
Used in large, unpressurized storage tanks where the entire full-to empty range must be measured.

FIG. 8-6. Chain and Tape Float Gauge

8.3.2.3: Lever and shaft float gauges (Fig. 8-7)
Used on either unpressurized or pressurized vessels where only a small range of level must be measured. The range of measurement is determined by the length of the float arm, but usually is between a few inches and a few feet.

FIG. 8-7. Lever and Shaft Float Gauge
8.3.2.4: Displacer level measuring device (Fig. 8-8)
One of the most frequently used level measuring devices is the torque tube displacer. It is attached to the free end of a torque tube which has elastic properties that permit it to twist as the displacer tries to float. This slight turning of the free end of the torque tube is connected to an indicator or transmitter.
Torque tube displacement gauges are normally limited to level spans of ten feet.

Fig. 8-8. Displacer Level Measuring Device

8.3.2.5: Head-pressure level gauges (Fig. 4-13)
The true level of a liquid can be determined by dividing the measured hydrostatic head by the density of the liquid. This method requires a knowledge of the densities of all phases of the liquid. Some of these methods are: pressure gauge, bubble tube, and differential pressure measurement. The bubbler (Fig. 4-13a) is used at vacuum and low pressures and is especially good for services such as molten sulfur and dirty liquids. In "boiling-liquid" service (Fig. 4-13b), a condensate trap must be used on the vapor leg. The level of trapped condensate in the vapor leg will usually be different than the vessel liquid level, requiring compensation of the transmitter.

8.3.2.6: Electrical type level gauges and switches (Fig.8-10)
Two common types of level gauges are the float-magnetic gauge configuration and the conductive type shown in Fig. 8-10. Slight tension on the tape reel permits the follower magnet to track the float at the liquid level in the device in Fig. 8-10a. The position of the reel represents the level and is either connected to an indicating device or a transmitter. The device shown in Fig. 8-10b illustrates the use of a conductive fluid for high and low level alarm indication.

Fig. 8-9. Head Pressure Level Gauges

Fig. 8-10. Electrical Level Gauges/Switches

8.3.2.7: Capacitance probes
A continuous method of level measurement based on electrical properties.
This method uses an electrode placed inside a vessel (or in a protective shell inside the vessel). The capacitance between the electrode and the wall of the vessel or shell varies as the dielectric constant varies. The dielectric in this case is the fluid, hence the capacitance varies in proportion to the liquid level. This capacitance is then measured, and converted to a level measurement to be indicated or transmitted.(figure 8-11)
These devices operate on the fact that process fluids generally have dielectric Oils have dielectric constants from 1.8 to 5. constants, , significantly different from that of air, which is very close to 1.0.
Pure glycol is 37; aqueous solutions are between 50 and 80. This technology requires a change in capacitance that varies with the liquid level, created by either an insulated rod attached to the transmitter and the process fluid, or an uninsulated rod attached to the transmitter and either the vessel wall or a reference probe. As the fluid level rises and fills more of the space between the plates, the overall capacitance rises proportionately. An electronic circuit called a capacitance bridge measures the overall capacitance and provides a continuous level measurement.

Figure 8-11. Capacitive level sensors measure the change in capacitance between two plates produced by changes in level. Two versions are available, one for fluids with high dielectric constants (A) and another for those with low dielectric constants (B).

8.3.2.8: Ultrasonic Level Transmitters.
Ultrasonic level sensors (see Figure 8-12) measure the distance between the transducer and the surface using the time required for an ultrasound pulse to travel from a transducer to the fluid surface and back. These sensors use frequencies in the tens of kilohertz range; transit times are ~6 ms/m. The speed of sound (340 m/s in air at 15°C (1115 fps at 60°F) depends on the mixture of gases in the headspace and their temperature. While the sensor temperature is compensated for (assuming that the sensor is at the same temperature as the air in the headspace), this technology is limited to atmospheric pressure measurements in air or nitrogen.
8.3.2.9: Radar Level Transmitters.
Through-air radar systems beam microwaves downward from either a horn or a rod antenna at the top of a vessel. The signal reflects off the fluid surface back to the antenna, and a timing circuit calculates the distance to the fluid level by measuring the round-trip time. these systems can be installed either vertically, or in some cases horizontally with the guide being bent up to 90° or angled, and provide a clear measurement signal.

Figure 8-12. Ultrasonic level measurement system, and Guided wave radar (GWR) systems
8.3.3: Temperature Sensors
8.3.3.1: Thermocouples
An ordinary thermocouple consists of two different kinds of wires (dissimilar metals) joined together at one end to form the measuring or “hot” junction. Where the free ends are connected to the measuring instrument, a reference or “cold” junction is formed. The millivolt readings measured by the instrument represent the difference in the temperatures of the two junctions and can be converted to temperature by various methods using conversion data from thermocouple tables. The reference temperatures normally used to generate thermocouple tables are 32°F and 70°F. Figs.8-13, 8-14.

Fig.8-13. Schematic drawing of a thermocouple

Fig.8-14. Schematic drawing of a thermocouple

Table. 8-2 shows some of the common thermocouple types, their usable temperature ranges, and the materials of construction.
Thermocouples used for process measurements are usually protected by a thermowell. The mass of the thermowell should be kept to a minimum in the interest of faster response. The thermocouple must be in thermal contact with the thermowell.
This is accomplished by the use of a thermally conductive lubricant or physical contact between the thermocouple and the well. In many measurement and control applications, electrical grounding of the thermocouple at the measurement point must be avoided.
Various series arrangements of thermocouples may be made to obtain differential temperatures or temperature averages.
Qualified personnel may check indicating or recording temperature devices measuring thermocouple potentials using portable equipment compatible with the thermocouple and with compensating circuitry identical to the primary device.
The use of incompatible equipment could result in erroneous results, especially in low temperature applications. At low temperatures, extreme care must be taken to eliminate sources of moisture in thermocouple installations. Common properties for different types of thermocouples are given in Table. 8-2. Conversion tables for converting millivolts to temperatures can be found in NBS Circular #561, or obtained from thermocouple suppliers for common types.

Table. 8-2. Thermocouple types.

8.3.3.2: Resistance thermometers
These are often called RTD’s for “Resistance Temperature Devices.” Since the resistance of metals changes as the temperature changes, a resistance thermometer can be constructed using this principle.
The metals that fit a near linear resistance temperature relationship requirement best are platinum, copper, and nickel. An accurate resistance measuring device utilizing a Wheatstone bridge is calibrated in units of temperature rather than resistance.
RTD’s are used in applications where faster responses and greater accuracies are required than may be obtained with thermocouples. Also RTD’s have a fairly high electrical output which is suitable for direct connection to indicators, controllers, recorders, etc. The use of RTD’s may also be more economical in some installations since the extension wires may be of copper rather than the more expensive thermocouple extension wire. A reference temperature source is not required for calibration. A special class of resistance thermometer is the thermistor device. It is low in cost, has fast response, and is very stable, but is limited to use at temperatures below 600°F.

8.3.3.3: Filled-system thermometers
These are simple, reliable, low cost devices. A bulb is attached to a capillary tube which is connected to a measuring element (bellows, Bourdon tube, etc.) in an indicating or transmitting device. The system is filled with a liquid or gas which changes in volume or pressure as the temperature of the bulb changes.

8.3.3.4: Glass stem thermometers
These devices are normally used in the office, laboratory, or other non-process areas.
Breakage is a problem; accuracy is from 0.1 to 2.0 degrees depending upon the range.

8.3.3.5: Bimetallic thermometers
The sensing element consists of two metals with different coefficients of expansion bonded together and attached to an indicator. These are inexpensive, but not very accurate and are normally used in on-off temperature thermostats where precise control is not required, or in process applications where relative changes are to be monitored. They should be calibrated at or near the normal operating point of the temperature being monitored.
8.3.4: Flow Sensors
8.3.4.1: Variable head flow meters
Flow meters in this class detect a pressure difference across a flow element specially designed to create that pressure difference. The most common flow element is the orifice plate, but other elements also in use are flow nozzles, venturi tubes, pitot tubes, averaging pitot tubes, target plates, and pipe elbows.

8.3.4.2: Variable area flow meters (Fig. 8-15)
This type includes the familiar rotameter. The differential pressure across the device is held constant, and the area through which the fluid passes changes due to the movement of the float up and down the tapered tube. These are usually limited to use with relatively small flows where visual indication is sufficient.

Fig.8-15. Rotameter

8.3.4.3: Turbine meters
These use a small permanent magnet mounted on the meter tube to create a magnetic field. A small turbine is mounted inside the tube and turns with a speed proportional to the flow rate. As each vane of the turbine passes through the magnetic field the magnetic flux is disturbed which induces a pulse in a pickup coil mounted on the outside of the meter. The pulse rate is proportional to the flow rate. Pulses are then counted and converted to standard flow units.

8.3.4.4: Positive displacement meters
Positive displacement meters and metering pumps measure discrete quantities of the flowing fluid. The rotating element is mechanically coupled to a transmitter or counter which integrates or totals the counts to provide an indication in units of gallons, liters, cubic feet, etc. Some common types are: rotating vane, bi-rotor, rotating paddle, oscillating piston, and oval gear meters. They are used for custody transfer devices such as gas meters or gasoline pumps.

8.3.4.5: Electromagnetic Flowmeter
If an electrical conductor is moved in a magnetic field, an electrical voltage is introduced in the conductor which is perpendicular to both the direction of motion and the magnet field and whose magnitude is proportional to the magnetic field strength and the velocity of the movement. The characterization of the laws of induction also applies to the movement of a conductive fluid in a pipe through a magnetic field and is the basis for the electrostatic flowmeter.
8.3.4.6: Ultrasonic Flow Meters
Ultrasonic flow measurement is based on sending and receiving acoustic signals through the flow. The difference in transit time between transducers, built in at opposite sides of the pipe gives signals that can be transferred to flow. A sound wave travels faster with the flow than one propagated against the flow. The difference in transit times is proportional to the medium’s mean flow velocity.
By installing more than one pair of transducers, a larger range of the flow profiles across the metering section can be covered and thereby increase the accuracy of the meter.

8.3.4.7: Other flowmeters
Some other flowmeter types occasionally encountered are:
• Doppler effect or ultrasonic flowmeters
• Vortex shedding flowmeters
• Laser velocimeters
• Thermal meters
• Nuclear Magnetic Resonance meters
• Gas ionization meters
• Cross-correlation devices
• Mass flowmeters
All flow meters should be calibrated using the fluid being measured, or, if a different fluid is used for calibration, the properties of the calibrating fluid must be related to the fluid of measurement.

8.4: Signal Transmitters
8.4.1: Pneumatic Transmitters
A pneumatic transmitter is a device that senses some process variable and translates the measured value into an air pressure which is transmitted to various receiver devices for indication, recording, alarm, and control. The signal range of 3-15 psig is the accepted industry standard; however, other ranges may be encountered. This signal is proportional to the range of measurement of the process variable. For example, 3-15 psig can represent 0-100 psi, 500-1000 gpm, –50 to +50°F, etc.
The prime function of a transmitter is to reproduce the low energy measurement signal with sufficient energy that it may be transmitted over an appreciable distance or used as a power source to a control device. The low-energy measurement signal is that position or movement associated with the action of the process variable on the sensing element (bellows, diaphragm, Bourdon tube, etc.). Pneumatic transmitters operate in a manner similar to proportional controllers.
8.4.2: Electronic Transmitters
Electronic transmitters perform the same function as pneumatic transmitters: a low energy process-related signal is converted into a higher energy signal suitable to connect to other instruments in the system. The output signal of most electronic transmitters is a 4-20 mA, 10-50 mA, or 1-5 Vdc signal.
Other ranges often encountered are: 0-10 Vdc, 2-10 Vdc, and 0.25-1.25 Vdc.
8.5: Signal Converters
Signal converters are used either to achieve compatibility between different types of instruments or for isolation purposes.
Some common forms of signal converters are:
8.5.1: Pneumatic-to-electronic (P/I)
These are electronic pressure transmitters designed for 3-15 psig input range and the desired output range (4-20 ma, etc.).
8.5.2: Electronic-to-pneumatic (I/P)
I/P converters are pneumatic transmitters with an electro-magnetic device connected to a nozzle-baffle arrangement which generates a pneumatic output signal which is proportional to the input signal.
8.5.3: Isolators
These are usually electronic current-to-current or voltage-to-voltage converters which provide electrical isolation to eliminate unwanted ground loop currents or common mode voltages.
8.5.4: Electric signal converters
These fit the same category as I/Ps and P/Is in that they change the signal from one range to another. Examples are 4-20 mA to 0-10 vdc, 1-5 Vdc to 10-50 mA, etc.
8.5.5: Frequency converters
Frequency to DC converters typically receive pulse inputs from turbines or positive displacement flowmeters and provide a proportional 4-20 ma, 10-50 mA or voltage output. Voltage output converters are often referred to as F/V (frequency-to-voltage) converters or transmitters.
V/F (voltage-to-frequency) converters are often used to interface standard “current-loop” type instrumentation to control devices requiring frequency or pulse-train set point inputs. These are commonly used in speed indicators for high speed centrifugal equipment.

8.6: Recorders and Indicators
8.6.1: Recorders
A recorder is a device used to plot the value of one or more measured variables, generally against time, but in some cases against another associated variable or variables. Recorders are often classified in the following ways:
1. According to use, i.e., whether the recorder is an integral part of the measuring/controlling system or is a general purpose type such as would be used in a laboratory or with a chromatograph.
2. According to method used to drive the pen(s). This refers to whether the pen is directly connected to the sensing element or to some type of pen positioning mechanism activated by the measuring signal.
3. According to chart type. This primarily refers to whether the recorder is of the circular or strip chart type and whether the time-axis drive is powered by a mechanical spring, electrical motor, or pneumatic drive.
4. Analog or Digital. Analog recorders are the more familiar strip chart and circular types. Digital recorders include such things as strip printers, data loggers, electronic totalizers, and computer-related devices such as data terminals and printers.
8.6.2: Indicators
An indicator is any device which presents a visual display of a measured quantity such as temperature, pressure, humidity, voltage, etc. Indicators are included in an instrumentation system either as independent devices (denoted as TI, PI, FI, etc.), or as a part of a controlling device (TIC, PIC, etc.). Indicators may be classified in the following groups:

8.6.2.1: Mechanical type
In these indicators the measured quantity causes the movement of a pointer along a graduated scale. This movement is due to the action of the measured quantity on a diaphragm, bellows, electromagnetic coil, or other sensing device which is mechanically linked to the pointer. This includes pressure gauges, filled tube dial thermometers, voltage and current meters, level gauges, etc.

8.6.2.2: Electronic analog type
These are analog indicators with no moving parts. A signal from the sensing device activates an optical display attached to graduated scale. A common type uses a bank of 200 tiny gas filled tubes which are illuminated additively in proportion to the magnitude of the process signal.

8.6.2.3: Digital type
Digital indicators include an analog-to digital converter which changes the electrical process signal to binary format which is then displayed in numerical form.
Typical displays consist of light emitting diodes (LED’s), liquid crystal displays (LCD’s), gas filled tubes, etc.

8.7: Control Concepts
8.7.1: Control Loops
A control circuit is commonly referred to as a “loop.” A control loop may be classified as either “open” or “closed” depending upon whether the control adjustments are manual settings (open loop) or automatically determined by some type of feedback controller (closed-loop).
8.7.1.1: Open loop (Fig. 8-16a)
In an open-loop control system, an operator makes a manual adjustment to a device (valve) which controls the flow of a manipulated variable (steam) to attempt to achieve some set-point (desired temperature) value of a controlled variable (hot water). However, this adjustment is only valid for the conditions under which the operator made the adjustment. Any disturbance such as a change in inlet water temperature, steam temperature, heat loss to the surroundings, or throughput will cause the outlet temperature to change.

Fig.8-16.a. Open loop.
8.7.1.2: Closed loop (Fig. 8-16b)
If appropriate measuring and controlling elements are added to the system, the loop is closed by the inclusion of an automatic feedback controller. The controller detects any difference between the set-point and measurement signals (error signal) and produces an output signal to drive the valve in the proper direction to adjust the heat input to cause the measurement to reach the set-point value.

Fig.8-16.b. closed loop.

8.7.1.3: Feedback control (Fig. 8-16c)
The basic components of a feedback control loop are shown in block diagram form in the figure. The “comparator” actually represents the entire controller and any associated signal converters. The “control element” is the valve, the “feedback element” is the transmitter, and the “process” is the mixing of the steam and cold water inside the water heater.

Fig.8-16.c. feedback control loop

8.7.1.4: Feedforward control (Fig. 8-16d)
Feed forward control (often called “Predictive Control”) is actually a form of open-loop control. An input variable (cold water temperature) is monitored and the manipulated variable (steam flow) is adjusted accordingly to compensate for changes in the input variable.
Feedforward control is almost always used in conjunction with feedback control to overcome the effects of some expected disturbance.

Fig.8-16.d. feedforward control loop

8.8: Control Modes and Controllers
Basic forms of control action or “modes” used in most process control are: two-position or “on-off” control, proportional control, integral or “reset” control, and derivative or “rate” control.
The latter three modes are often used in various combinations with each other.

8.8.1: Two-Position (on-off) Controllers
The simplest form of control action is “on-off” control, in which the controller output either energizes or de-energizes some two-state device such as a relay or an open-shut type valve. The two-position controller is used extensively in home heating and cooling systems, refrigerators, hot water tanks, air compressors, and other applications where the cost of more precise control is not justified. Most two-position controllers are reverse-acting, i.e., when the measured variable is above the set-point, the controller turns the manipulated variable OFF, and when the measured variable is below the set point, the controller turns the manipulated variable ON. A “deadband” or differential gap exists around the zero error condition to minimize cycling. This is often implemented as a pair of control points: one where the controller will “kick-on” and the other where the controller will “kick-off” as opposed to a single set point. Fig.8-17.
8.8.2: Proportional Control Mode
In the proportional control mode, the final control element is throttled to various positions that are dependent on the process system conditions. For example, a proportional controller provides a linear stepless output that can position a valve at intermediate positions, as well as "full open" or "full shut."

Fig.8-17. two positions control system.

Fig.8-18. Proportional system control system.

8.9: Control Valves
Selecting the proper control valve for each application involves many factors. The valve body design, actuator style, and plug characteristic are critical items for selection. Proper valve sizing is necessary for accurate, efficient, economical process control. In areas where personnel will be affected, noise prediction and control becomes a significant factor.
Engineering application guidelines, nomographs, and equations presented in the following pages may be used to determine the correct control valve configuration, size and flow characteristics, and to predict noise levels for most applications.
The material presented here may also be used to evaluate the performance of valves installed in existing plants. The equations given in this section are used to calculate the flow coefficient (Cv or Cg) required for a valve to pass the required flow. Most valve manufacturers publish flow coefficients for each valve style and size.
A brief description of the two major components of a control valve, the valve body and the actuator, is presented in Fig. 8-19.

FIG. 8-19. Relationship of Major Components

8.9.1: Control-Valve Bodies
The control-valve body (see Fig. 8-20) regulates the rate of fluid flow as the position of the valve plug is changed by force from the actuator. Therefore, the valve body must permit actuator thrust transmission, resist chemical and physical effects of the process, and provide the appropriate end connections to mate with the adjacent piping. It must do all of this without external leakage. Most valve body designs are of the globe style, but other configurations such as ball and butterfly styles are available. Final selection depends upon detailed review of the engineering application.

FIG. 8-20. Control valve body.

8.9.2: Control-Valve Actuators
Pneumatically operated control-valve actuators are the most popular type in use, but electric, hydraulic, and manual actuators are also widely used. The spring-and-diaphragm pneumatic actuator (see Figs. 8-21a and b.) is commonly specified, due to its dependability and its simplicity of design. Pneumatically operated piston actuators provide integral positioner capability and high stem-force output for demanding service conditions, such as high differential pressure or long valve stem travel distance.

FIG. 8-21A . Direct Acting Spring-and-Diaphragm Actuator Assemblies (Air to close, fail open)

FIG. 8-21b. Typical Reverse Acting Spring-and-Diaphragm Actuator Assemblies. (Air to open, fail close)

8.9.3: Flow Characteristics and Valve Selection
The flow characteristic of a control valve is the relationship between the flow rate through the valve and the valve travel as the travel is varied from 0 to 100%.
Fig. 8-20 illustrates typical flow-characteristic curves.
• The quick-opening flow characteristic provides for maximum change in flow rate at low valve travel with a fairly linear relationship. Additional increases in valve travel give sharply reduced changes in flow rate. When the valve plug nears the wide open position, the change in flow rate approaches zero.
In a control valve, the quick-opening valve plug is used primarily for on-off service; however, it is also suitable for many applications where a linear valve plug would normally be specified.
• The linear flow-characteristic curve shows that the flow rate is directly proportional to the valve travel. This proportional relationship produces a characteristic with a constant slope so that with constant pressure drop (ΔP), the valve gain will be the same at all flows. (Valve gain is the ratio of an incremental change in flow rate to an incremental change in valve plug position. Gain is a function of valve size and configuration, system operating conditions, and valve plug characteristic.)
The linear-valve plug is commonly specified for liquid level control and for certain flow control applications requiring constant gain.
• In the equal-percentage flow characteristic, equal increments of valve travel produce equal percentage changes in the existing flow. The change in flow rate is always proportional to the flow rate just before the change in position is made for a valve plug, disc, or ball position. When the valve plug, disc, or ball is near its seat and the flow is small, the change in flow rate will be small; with a large flow, the change in flow rate will be large. Valves with an equal-percentage flow characteristic are generally used for pressure control applications.
They are also used for other applications where a large percentage of the total system pressure drop is normally absorbed by the system itself, with only a relatively small percentage by the control valve. Valves with an equal-percentage characteristic should also be considered where highly varying pressure drop conditions can be expected.
• The modified parabolic-flow characteristic curve falls between the linear and the equal-percentage curve.

Note: Where detailed process knowledge is lacking, as a rule of thumb, use equal-percentage characteristics at 70% opening for the valve sizing.

FIG. 8-22. Example Flow Characteristic Curves.

8.9.4: Fundamentals of Control Valve Sizing
8.9.4.1: Gas Service
Critical Pressure Drop — Critical flow limitation is a significant problem with sizing valves for gaseous service.
Critical flow is a choked flow condition caused by increasing gas velocity at the vena contracta. The vena contracta is the point of minimum cross-sectional area of the flow stream which occurs just downstream of the actual physical restriction.
When the velocity at the vena contracta reaches sonic velocity, additional increases in pressure drop, ΔP, (by reducing downstream pressure) produce no increase in flow.
In the ISA sizing procedure critical flow limitations i addressed by calculating (Υ), the expansion factor, for utilization within the actual sizing equation.

Υ = 1 – X / (3FkXc) Eqn. 8-1
Where
Fk = k /1.4 Eqn. 8-2
Where
Υ = expansion factor, ratio of flow coefficient for a gas to that for a liquid at the same Reynolds number, dimensionless
X = ratio of pressure drop to absolute inlet pressure (ΔP/P1), dimensionless
Fk = ratio of specific heats factor, dimensionless, for natural gas use k = 1.27, so Fk = 0.9071
Xc = pressure drop ratio for the subject valve at critical flow, with Fk = 1.0, dimensionless (from table 8-3)
k = ratio of specific heats, dimensionless. For natural gas use k = 1.27

Critical pressure drop, and thus critical flow, is realized when X ≥ FkXc. Therefore, since the flow can’t exceed that produced at the critical pressure drop the value of Υ in the following sizing equations should never be less than 0.67.

Υ = 1 – X / (3FkXc) = 1- (1/3) = 0.67 Eq 8-3

Likewise the value of X in the equations should never exceed FkXc.

Sizing Calculation Procedure – The compressible fluid sizing equation can be used to determine the flow of gas or vapor through any style of valve. Absolute units of temperature and pressure must be used in the equation.
Equations used to calculate the required Cv (valve flow coefficient) and thus valve size for a given set of service conditions is as follows:.

Q = 32640 FpCvP1Y [X/(GgTZ)]0.5 Eq.8-4

Where
Q = volumetric flow rate scfd
Fp=piping geometry factor, dimensionless (If the valve has no such fittings attached, e.g., the nominal valve size and nominal pipe size are the same, then Fp = 1.0), for other configuration Fp value can be calculated per the ANSI/ISA S75.01 standard.
Cv = valve flow coefficient. Given from valve data sheets or to be calculated using known flow rate for valve design.
P1 = upstream absolute static pressure, measured two nominal pipe diameters upstream of valve fitting assembly, psia
Y = expansion factor, calculated from eq. 8-1 and 8-2 and should never be less than 0.67
X = (ΔP/P1)
Gg = Gas specific gravity.
T = Temperature 0R.
Z = Compressibility factor.

The equations can likewise be rearranged to calculate the flow or pressure drop for a given valve and set of service conditions.
For a new valve selection, a valve size is typically chosen such that the maximum, calculated Cv is close to 75% to 85% of valve travel. This allows for process variability while maintaining flow capability. The minimum, calculated Cv should typically occur at or about 10% of valve travel.

8.9.4.2: Liquid Service
The procedure used to size control valves for liquid service should consider the possibility of cavitation and flashing since they can limit the capacity and produce physical damage to the valve. In order to understand the problems more thoroughly, a brief discussion of the cavitation and flashing process is presented below.

Flow Capacity --- The valve sizing coefficient most commonly used as a measure of the capacity of the body and trim of a control valve is Cv. One Cv is defined as one U.S. gallon per minute of 60 0F water that flows through a valve with a one psi pressure drop. The general equation for Cv is as follows:
Cv = flow x (sp.gr at flowing temperature/ΔP)0.5 Eq.8-5
When selecting a control valve for an application, the calculated Cv is used to determine the valve size and the trim size that will allow the valve to pass the desired flow rate and provide stable control of the process fluid.

Table 8-3. Typical Cv, Xc and FL Values for Valves

Cavitation — In liquids, when the pressure anywhere in the liquid drops below the vapor pressure of the fluid, vapor bubbles begin to form in the fluid stream. As the fluid decelerates there is a resultant increase in pressure. If this pressure is higher than the vapor pressure, the bubbles collapse (or implode) as the vapor returns to the liquid phase.
Cavitation occurs in pumps, control valves, and any flow chocking devices.
Determining when a problem-causing level of cavitation is present represents a considerable challenge. The reader is referred to ISA RP75.23, “Considerations for Evaluating Control Valve Cavitation.” This recommended practice provides more information on the cavitation process as well as suggesting a common terminology and methodology for making safe valve selections in cavitating applications.
As discussed in the recommended practice, the selection of the appropriate operating limit for a given situation is dependent on the service conditions but should also consider other influences such as duty cycle, location, desired life, and past experience.
All of these point to the need to consult the valve manufacturer when selecting a valve for cavitation control.

Fig. 8-23. Bubble formations.

Fig. 8-24. Bubble compression and collision with metal surface.

Flashing — If the downstream pressure is equal to or less than the vapor pressure, the vapor bubbles created at the vena contracta do not collapse, resulting in a liquid-gas mixture downstream of the valve. This is commonly called flashing. When flashing of a liquid occurs, the inlet fluid is 100 percent liquid which experiences pressures in and downstream of the control valve which are at or below vapor pressure. The result is a two phase mixture (vapor and liquid) at the valve outlet and in the downstream piping. Velocity of this two phase flow is usually very high and results in the possibility for erosion of the valve and piping components.

Sizing Information — The following section is based on ISA-S75.01, “Flow Equations for Sizing Control Valves.” The reader is referred to that standard for more complete discussion of these equations and methods. As that standard points out, these equations are not intended for situations involving mixed-phase fluids, dense slurries, dry solids, or non-Newtonian liquids. In these cases the valve manufacturer should be consulted for sizing assistance.
The ISA methodology recognizes the impact of service conditions that will cause the liquid to vaporize at some point between the inlet and outlet of the valve. This vaporization results in either cavitation or flashing, causing a breakdown in the normal relationship between Cv and √ΔP and ultimately a limit to the flow through the valve regardless of an increasing pressure drop caused by decreasing P2. The recognition of this comes in the form of a separate sizing equation for each regime, nonvaporizing and vaporizing. Each must be solved and then the larger calculated Cv chosen as the required value.
This discussion of liquid sizing will be further restricted to:
1. Turbulent flow streams: There are usually flow streams that are not either high viscosity or low velocity. The majority of process plant control valves do operate in the turbulent regime, however if the Reynolds number for a process is less than 4000 the reader is referred to the ISA standard where a non-turbulent flow correction method can be found.
2. Valve installed without fittings attached to the valve ends: When fittings are present there are, as with the previous gas sizing discussion, necessary modifications to the sizing equations to accommodate the additional disturbance to flow. This discussion will be limited to the case where there are no fittings attached, therefore the valve size and pipe size are the same, Fp = 1.0. Refer to the full ISA standard for the proper methods if fittings are present.

Sizing Calculation Procedure —

Nonvaporizing Volumetric flow (gpm) q = FpCv (ΔP/Gf)0.5 Eq. 8-6

Vaporizing Volumetric flow (gpm) q = FLCv [(P1-FFPv)/Gf]0.5 Eq. 8-7

Where
q = Volumetric flow (gpm)
Fp = piping geometry factor, dimensionless (If the valve has no such fittings attached, e.g., the nominal valve size and nominal pipe size are the same, then Fp = 1.0), for other configuration Fp value can be calculated per the ANSI/ISA S75.01 standard.
Cv = valve flow coefficient. Given from valve data sheets or to be calculated using known flow rate for valve design.
P1 = upstream absolute static pressure, measured two nominal pipe diameters upstream of valve fitting assembly, psia
FF = liquid critical pressure ratio factor, dimensionless. From fig. 8-25 based on the critical pressure and inlet vapor pressure for subject liquid using Table. 8-4, which lists critical pressures for some common fluids.
FL = liquid pressure recovery factor of a valve without attached fittings, dimensionless. Table 8-3.
Pv = vapor pressure of liquid at valve inlet temperature, psia
Gf = liquid specific gravity at upstream conditions, dimensionless

Fig. 8-25. Liquid critical pressure ratio factor, FF

Table. 8-4. Critical pressure for selected liquids, psia.

yasserkassem · Post by **yasserkassem** » Sun Mar 22, 2020 1:39 pm

Chapter 9

Pressure Relief System

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Chapter 9 256
Process Relief Systems 256
9.1: Introduction 256
9.2: Relief Device Design and Requirements: 256
9.2.1: Blocked Discharge 257
9.2.2: Fire Exposure 257
9.2.3: Tube Rupture 257
9.2.4: Control Valve Failure 257
9.2.5: Thermal Expansion 257
9.2.6: Utility Failure 257
9.3: General discussion 258
9.4: Special Relief System Considerations 260
9.4.1: Pumps and storage equipment 260
9.4.2: Low Temperature Flaring 260
9.5: Relieving Devices 260
9.5.1: Conventional Relief Valves 260
9.5.2: Balanced Relief Valves 262
9.5.3: Pilot Operated Relief Valves 262
9.5.4: Resilient Seat Relief Valves 264
9.5.5: Rupture Disk 265

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Chapter 9

Process Relief Systems

9.1: Introduction
The most important safety devices in a production facility are the pressure relief valves, which ensure that pipes, valves, fittings, and pressure vessels can never be subjected to pressures higher than their design pressures.
Relief valves must be designed to open rapidly and fully, and be adequately sized to handle the total flow of gas and liquids that could potentially cause an overpressure situation. They relieve the pressure by routing this stream to a safe location where it can be vented to atmosphere or burned.
As long as pressure, level, and temperature control devices are operating correctly, the safety system is not needed “In case of steady conditions and working within design range”. If the control system malfunctions, then pressure, level, and temperature safety switches sense the problem so the inflow can be shut off. If the control system fails and the safety switches don't work, then relief valves are needed to protect against overpressure. Relief valves are essential because safety switches do fail or can be bypassed for operational reasons. Also, even when safety switches operate correctly, shutdown valves take time to operate, and there may be pressure stored in upstream vessels that can overpressure downstream equipment while the system is shutting down. Relief valves are an essential element in the facility safety system.

9.2: Relief Device Design and Requirements:
The ASME code requires every pressure vessel that can be blocked in to have a relief valve to alleviate pressure build up due to thermal expansion of trapped gases or liquids. In addition, the American Petroleum Institute Recommended Practice (API RP) 14C, "Analysis, Design, Installation and Testing of Basic Surface Safety Systems on Offshore Production Platforms," recommends that relief valves be installed at various locations in the production system; and API RP 520, "Design and Installation of Pressure Relieving Systems in Refineries," recommends various conditions for sizing relief valves.
Proper selection, use, location, and maintenance of relief devices are essential to protect personnel and equipment as well as to comply with codes and laws.
Determination of the maximum relief required may be difficult.
Loads for complex systems are determined by conservative assumptions and detailed analysis. By general assumption, two unrelated emergency conditions caused by unrelated equipment failures or operator error will not occur simultaneously (no double jeopardy). The sequence of events must be considered. The development of relief loads requires the engineer to be familiar with overall process design, including the type of pump drives used, cooling water source, spares provided, plant layout, instrumentation, and emergency shutdown philosophy.
In production facility design, the most common relieving conditions are as follows:

9.2.1: Blocked Discharge
The outlet of almost any vessel, pump, compressor, fired heater, or other equipment item can be blocked by mechanical failure or human error. In this case, the relief load is usually the maximum flow which the pump, compressor, or other flow source produces at relief conditions.

9.2.2: Fire Exposure
Fire is one of the least predictable events which may occur in a gas processing facility, but is a condition that may create the greatest relieving requirements. If fire can occur on a plant-wide basis, this condition may dictate the sizing of the entire relief system; however, since equipment may be dispersed geographically, the effect of fire exposure on the relief system may be limited to a specific plot area. Vapor generation will be higher in any area which contains a large number of uninsulated vessels. Various empirical equations have been developed to determine relief loads from vessels exposed to fire. Formula selection varies with the system and fluid considered. Fire conditions may overpressure vapor-filled, liquid-filled, or mixed-phase systems.

9.2.3: Tube Rupture
When a large difference exists between the design pressure of the shell and tube sides of an exchanger (usually a ratio of 1.5 to 1 or greater), provisions are required for relieving the low pressure side. Normally, for design, only one tube is considered to rupture. Relief volume for one tube rupture can be calculated using appropriate sizing equations in this section.
When a cool media contacts a hot stream, the effects of flashing should be considered. Also the possibility of a transient overpressure caused by the sudden release of vapor into an all-liquid system should be considered.
9.2.4: Control Valve Failure
The failure positions of instruments and control valves must be carefully evaluated. In practice, the control valve may not fail in the desired position. A valve may stick in the wrong position, or a control loop may fail. Relief protection for these factors must be provided. Relief valve sizing requirements for these conditions should be based on flow coefficients (manufacturer data) and pressure differentials for the specific control valves and the facility involved.
9.2.5: Thermal Expansion
If isolation of a process line on the cold side of an exchanger can result in excess pressure due to heat input from the warm side, then the line or cold side of the exchanger should be protected by a relief valve.
If any equipment item or line can be isolated while full of liquid, a relief valve should be provided for thermal expansion of the contained liquid. Low process temperatures, solar radiation, or changes in atmospheric temperature can necessitate thermal protection. Flashing across the relief valve needs to be considered.
9.2.6: Utility Failure
Loss of cooling water may occur on an area-wide or plant-wide basis. Affected are fractionating columns and other equipment utilizing water cooling. Cooling water failure is often the governing case in sizing flare systems. Electric power failure, similar to cooling water failure, may occur on an area-wide or plant-wide basis and may have a variety of effects. Since electric pump and air cooler fan drives are often employed in process units, a power failure may cause the immediate loss of reflux to fractionators. Motor driven compressors will also shut down. Power failures may result in major relief loads.
Instrument air system failure, whether related to electric power failure or not, must be considered in sizing of the flare system since pneumatic control loops will be interrupted. Also control valves will assume the position as specified on "loss of air" and the resulting effect on the flare system must be considered.
A vessel may be subject to more than one condition under different failure scenarios.
For example, a low pressure separator may be subject to blocked discharge, gas blowby from the high pressure separator, and fire. Only one of these failures is assumed to happen at any time. The relief valve size needs to be calculated for each pertinent relieving rate and the largest size used. The usual controlling cases for common vessels and piping are shown in Table 9-1.

Vessel Relieving Scenario
Production Separators Blocked Discharge
Test Separators Blocked Discharge
Low Pressure Separators Blocked Discharge or Blowby
Glycol Contact Tower Fire
Oil Treater Gas Blowby or Fire
Utility or Fuel Gas Scrubber Regulator Failure
Heat Exchanger Tube Rupture
Compressor Scrubber Fire
Compressor Discharge Blocked Discharge

Table 9-1. Maximum Rate Relieving Scenarios
9.3: General discussion
A vessel can only be overpressurized if the upstream vessel has a higher pressure than the vessel in question. A compressor scrubber with a MAWP of 285 that gets flow from a 285 MAWP separator does not need to have a relief valve sized for blocked discharge. The upstream relief valve will keep the upstream separator pressure from going higher than 285, so there is no way it can overpressure the downstream scrubber. The scrubber PSV only needs to be sized for fire.

Good engineering judgment should be used to determine the relief rate when the separator MAWP is higher than the well SITP (shut in tubing pressure). Unexpected things can happen with a well. Production reservoirs at different pressures within the well bore can communicate in unexpected ways (for example, as the result of a poor cement job). Where flow is coming from a well, it is a good idea to provide an extra margin of safety. In the same time, the relief valve should be sized for blocked discharge of the full production rate.

Figure 9-1 shows the various relationships between MAWP and the relief valve set pressure. The primary relief valve should be set to open at no more than 100% of MAWP and to relieve the worst case flow rates, not counting fire (i.e., blocked discharge or gas blowby), at a pressure of 1.10 MAWP. If two relief valves are used to handle the worst case flow rates, the first must be set no higher than 100% MAWP and the second at 1.05 MAWP. They must relieve the worst case flow rates, not counting fire, at 1.16 MAWP. The maximum pressure for relieving fire relief rates is 1.21 MAWP. Thus, under relief conditions, the pressure in the vessel may actually exceed MAWP. This buildup of pressure in the vessel above the MAWP as the relief valve opens is called "overpressure." This is taken into account by the various safety factors in the ASME Code and is one of the reasons the vessel is originally tested to 1.5 MAWP.

The relief valve must be installed so that gases are routed to a safe location. In small facilities and remote locations this is accomplished with a simple "tail pipe," which points the discharge vertically upward and creates a jet in excess of 500 feet per second. The jet action dilutes the discharge gases to below the lower flammable limit in approximately 120 pipe diameters. Liquids may fall back on the equipment.
In large facilities and offshore platforms where the escaping gases and liquids could present a source of pollution or ignition, it is common to route the relief valve discharges into a common "header" that discharges at a remote safe location. Often a vent scrubber is installed in this header to separate the bulk of the liquids and to minimize the possibility of liquid discharges to atmosphere.

Pressure relief valve is a generic term applied to relief valves, safety valves, or safety relief valves. Definition by type of relief valve is covered in the relief device description. Relief valve characteristics related to pressure vessel requirements are illustrated in Fig. 9-1.

Fig. 9-1. Characteristics of Safety Relief Valves for Vessel Protection

Notes:
1. The operating pressure may be any lower pressure required.
2. The set pressure and all other values related to it may be moved downward if the operating pressure permits.
3. This figure conforms with the requirements of the ASME Boiler and Pressure Vessel Code, Section VIII.
4. The pressure conditions shown are for safety relief valves installed on a pressure vessel (vapor phase).
9.4: Special Relief System Considerations
9.4.1: Pumps and storage equipment
The following considerations should be followed for relief system design in the following equipment.
Pumps — Relief valves are required on the discharge of each positive displacement pump. Normally, these reliefs are piped back to the process upstream of the pump rather than to the flare system. Isolation valves around the relief valves may not be required if the pump itself can be isolated for maintenance.
Vessels and Tanks — Vessels or tanks which are subject to atmospheric "breathing" due to cooling of gas or liquid contents are normally protected by "breather" valves or vacuum relief valves.
Compressors — Each positive displacement compressor must have a relief valve on the discharge of each stage upstream of the block and check valves in order to protect the compressor.

9.4.2: Low Temperature Flaring
When low temperature streams are relieved, the flare system piping and equipment exposed to cryogenic temperature may require stainless steel or other acceptable alloys.
The system should be designed for the coldest process stream to be relieved plus the lower temperature effect of the expanding fluid (Joule-Thomson effect). Materials selection should be made according to applicable code recommendations.
9.5: Relieving Devices
Valves that activate automatically to relieve pressure are called "safety valves," "relief valves," or "safety relief valves."
Safety valves are spring loaded and characterized by a rapid full opening or "pop" action. They are used primarily for steam or air service. Sometimes they are referred to as "pop valves."
Relief valves are spring loaded and open more slowly. They reach full opening at 25% over set pressure and are used primarily for liquid services.
Safety relief valves can be either spring loaded or pilot operated and are designed to provide full opening with little overpressure.
Most automatically-actuating relief devices used in production facilities are actually safety relief valves; however, they are commonly referred to as relief valves or safety valves. In this book, the term "relief valve" is used in the generic sense of any automatically-actuating pressure relieving device.
There are three types of relief valves: conventional, balanced-bellows, and spring loaded.

9.5.1: Conventional Relief Valves
In a conventional relief valve, the inlet pressure to the valve is directly opposed by a spring. Tension on the spring is set to keep the valve shut at normal operating pressure but allow the valve to open when the pressure reaches relieving conditions.
This is a differential pressure valve. Most conventional safety-relief valves available to the petroleum industry have disks which have a greater area, AD, than the nozzle seat area, AN. The effect of back pressure on such valves is illustrated in Fig. 9-2. If the bonnet is vented to atmospheric pressure, the back pressure acts with the vessel pressure so as to overcome the spring force, FS, thus making the relieving pressure less than when set with atmospheric pressure on the outlet. However, if the spring bonnet is vented to the valve discharge rather than to the atmosphere, the back pressure acts with the spring pressure so as to increase the opening pressure. If the back pressure were constant, it could be taken into account in adjusting the set pressure. In operation the back pressure is not constant when a number of valves discharge into a manifold. A cut-away of a conventional relief valve is shown in Fig. 9-2.

FIG. 9-2.Conventional Safety-Relief Valve, and effect of back pressure on valve setting.
Conventional relief valves should only be used where the discharge is routed independently to atmosphere, or if installed in a header system, the back-pressure build-up when the device is relieving must be kept below 10% of the set pressure so the set point is not significantly affected. (The set point increases directly with back-pressure).
Conventional relief valves may be equipped with lifting levers or screwed caps. The lifting lever permits mechanical operations of the valve for testing or clean-out of foreign material from under the seat.

9.5.2: Balanced Relief Valves
Balanced safety-relief valves incorporate means for minimizing the effect of back pressure on the performance characteristics — opening pressure, closing pressure, lift, and relieving capacity.
These valves are of two types, the piston type and the bellows type, as shown diagrammatically in Fig. 9-3. In the piston type, of which several variations are manufactured, the guide is vented so that the back pressure on opposing faces of the valve disk cancels itself; the top face of the piston, which has the same area, AP, as the nozzle seat area, AN, is subjected to atmospheric pressure by venting the bonnet. The bonnet vented gases from balanced piston-type valves should be disposed of with a minimum restriction and in a safe manner.
In the bellows type of balanced valve, the effective bellows area, AB, is the same as the nozzle seat area, AN, and, by attachment to the valve body, excludes the back pressure from acting on the top side of that area of the disk. The disk area extending beyond the bellows and seat area cancel, so that there are no unbalanced forces under any downstream pressure.
The bellows covers the disk guide so as to exclude the working fluid from the bonnet. To provide for a possible bellows failure or leak, the bonnet must be vented separately from the discharge. The balanced safety-relief valve makes higher pressures in the relief discharge manifolds possible.
Both balanced-type valves shown in Fig. 9-3 should have bonnet vents large enough to assure no appreciable back pressure during design flow conditions. If the valve is in a location in which atmospheric venting (usually not a large amount) presents a hazard, the vent should be piped to a safe location. The user should obtain performance data on the specific type of valve being considered.
A cross section drawing of a balanced (bellows) relief valve is shown in Fig. 9-3.

9.5.3: Pilot Operated Relief Valves
A pilot operated relief valve consists of two principal parts, a main valve and a pilot. The valve utilizes a piston instead of a shaft. Inlet pressure is directed to the top of the main valve piston. More area is exposed to pressure on the top of the piston than on the bottom so pressure, instead of a spring, holds the main valve closed. At the set pressure, the pilot opens, reducing the pressure on top of the piston thereby allowing the main valve to open fully. For some applications, pilot-operated relief valves are available in minimum friction, light-weight diaphragm construction (in place of heavy pistons). Pilot operated valves can allow backflow if downstream pressure exceeds set points. Backflow prevention is required on valves connected to common relief headers. A check valve, split piston type valve, or backflow preventer in the pilot line can be used.

A typical pilot operated relief valve is shown in Fig 9-4. This style valve should be considered for applications involving high back pressure, high operating pressure, or where premium seat tightness is desired.

FIG. 9-3. Balanced Safety-Relief Valve and the effect of back pressure on set pressure.

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FIG. 9-4. Pilot Operated Relief Valve

9.5.4: Resilient Seat Relief Valves
With the use of metal-to-metal seat conventional or balanced type relief valves where the operating pressure is close to the set pressure, some leakage can be expected through the seats of the valve (refer to API Standard 527). Resilient seat relief valves with either an O-ring seat seal or plastic seats can provide seat integrities which exceed API Standard 527 (Fig. 9-5); however, there are limitations of temperature and material compatibility when using these valves, and manufacturer guidelines should be consulted. Although such valves provide near zero leakage until seat damage occurs, the resilient seats may erode rapidly once leakage begins.

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FIG. 9-5. O-Ring Seals - Conventional and Bellow Valves

9.5.5: Rupture Disk
A rupture disk consists of a thin diaphragm held between flanges. The disk is designed to rupture and relieve pressure within tolerances established by ASME Code. Rupture disks can be used in gas processing plants, upstream of relief valves, to reduce minor leakage and valve deterioration. In these installations, the pressure in the cavity between the rupture disk and the relief valve should be monitored to detect a ruptured disk. In some applications a rupture disk with a higher pressure rating is installed parallel to a relief valve. A rupture disk is subject to fatigue failure due to operating pressure cycles.
To establish recommended replacement intervals, consult rupture disk suppliers.
Rupture disks should be used as the primary relieving device only if using a pressure relief valve is not practical. Some examples of such situations are:
(a) Rapid rates of pressure rise. A pressure relief valve system does not react fast enough or cannot be made large enough to prevent overpressure, e.g., an exchanger ruptured tube case or a runaway reaction in a vessel.
(b) Large relieving area required. Because of extremely high flow rates and/or low relieving pressure, providing the required relieving area with a pressure relief valve system is not practical.
(c) A pressure relief valve system is susceptible to being plugged, and thus inoperable, during service.

Figure.9-6. Rupture disk.

Figure. 9-7. Relief valve and safety relief valve
Fundamentals of Oil and Gas Processing
References.
1- Gas Processors Suppliers Association GPSA Engineering Data Book 11th, 12th, & 15th Editions. Tusla, OK.
2- 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.
3- Abdel-Aal, H. K., Surface Petroleum Operations, Saudi Publishing & Distributing House, Jeddah, 1998.
4- H.K. Abdel-Aal and Mohamed Aggour, Petroleum and gas field processing, 2003 by Marcel Dekker, Inc.
5- Crude-Oil-Treating-Systems-Design-Manual-Sivalls-Inc.
6- API Spec. 12J (Specification of Oil and Gas Separator) 7th.ed. Oct. 1998.
7- API Spec. 12K (Indrict Fired Heater) 7th.ed.Jun.1989.
8- API Spec. 12L (Vertical and Horizental Emulsion Treater) 4th ed. Nov. 1994.
9- API Std. 560 (Fired Heaters for General Refinery Service) 3rd. ed. May 2001
10- Standard Handbook of Petroleum and Natural Gas 2nd ed. William C. Lyons, Ph.D., P.E. Gary J. Plisga, B.S. 2005, Elsevier Inc.
11- Gas Pipeline Hydraulics, E. Shashi Menon, P.E. PDH Engineering course material.
12- Chilingarian, G. V., Robertson, J. O., Jr., and Kumar, S., Surface Operations in Petroleum Production, I, 1987, Elsevier Science, Amsterdam.
13- The Chemistry and Technology of Petroleum, James G. Speight
14- Flow Management for Engineers and Scientists, Nicholas P. Cheremisinoff and Paul N. Cheremisinoff.
15- Al-Tahini, A., Crude oil Emulsions, Co-op Report, Department of Chemical Engineering, KFUPM, Dhahran, Saudi Arabia, 1996.
16- Thro, M. E. and Arnold, K. E., Water droplet size determination for improved oil treater sizing, SPE 69th Annual Technical Conference and Exhibition, 1994.
17- Basseler, O. U., De-emulsification of Enhanced Oil Recovery Produced Fluids, Tretolite Div., Petrolite Corp., St. Louis, MO, 1983.
18- Manning, F. S. and Thomson, R., Oil-Field Processing of Petroleum, Penn-well Publishing, Tulsa, OK, 1991.
19- Campbell, John M., ‘‘Gas Conditioning and Processing,’’ Vol. 2, published by Campbell Petroleum Series, Norman, Oklahoma, 1976.
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Fundamentals of Oil and Gas Processing Book "full text"

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