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Thread: Basics of Gas Field Processing Book "Full text"

  1. Re: Basics of Gas Field Processing Book "Full text"

    Gas Dehydration - Chapter 5 - Part 4 Regeneration Gas Separator
    Most desiccants also have an affinity for hydrocarbons, thus a skimmer is used to separate the valuable hydrocarbons from the water to be discarded. Frequent pH tests on the discarded water helps pinpoint corrosion problems in the adsorption system. A common problem encountered in regeneration gas separators is the fouling of the liquid dump line by desiccant dust and heavy oils. Expander Plant Molecular Sieve Applications
    Turbo expander plants commonly operate down to temperatures of -150 0F.
    Operating points much below the equilibrium water content data illustrated in McKetta-Wehe chart (include designs to water contents as low as 1 ppm).
    As shown in Table 5-7 (UP), only molecular sieves and activated alumina are capable of such performance. Molecular sieves are used in approximately 95% of the dehydration equipment for this type of plant (a 4A molecular sieve has twice the adsorptive capacity of activated alumina).
    5.23 Desiccant Performance
    5.23.1 General Conditions
    Desiccants decline in adsorptive capacity at different rates under varying operating conditions.
    Desiccant aging is a function of many factors, including:
    Number of cycles experienced
    Exposure to any harmful contaminants present in the inlet stream.
    The most important variable affecting the decline rate of desiccant capacity is the chemical composition of the gas or liquid to be dried.

    Capacity of a new desiccant will decline slowly during the first few months in service because of cyclic heating, cooling, and netting. Desiccant capacity usually stabilizes at about 55 to 70% of the initial capacity.

    5.23.2 Moisture Analyzer
    The moisture analyzer is used to optimize the drying cycle time, where the drying time is always shortened as the desiccant ages. Both inlet and outlet moisture analyzer probes should be used. A probe extending approximately 2 feet upward into the bed from the outlet end is recommended because it allows a dehydration capacity test to be run without the risk of a water breakthrough.

    5.23.3 Effect of Contaminants in Inlet Feed Stream
    Compressor oils, corrosion inhibitors, glycols, amines, and other high-boiling contaminants cause a decline in desiccant capacity, because normal reactivation temperatures will not vaporize the heavy materials. Residual contaminants slowly build up on the desiccants surface reducing the area available for adsorption. Many corrosion inhibitors chemically attack certain desiccants, permanently destroying their usefulness.

    5.23.4 Effect of Regeneration Gases Rich in Heavy Hydrocarbons
    Use of this rich gas in a 5500 to 600 0F regeneration service aggravates coking problems. Rich gases may be dried satisfactorily with molecular sieves. Lean dry gas is always preferable for regeneration.

    5.23.5 Effect of Methanol in the Inlet Gas Stream
    Methanol in the inlet gas is a major contributor to the coking of molecular sieves where regeneration is carried out at temperatures above 550 0F. Polymerization of methanol during regeneration produces dimethyl ether and other intermediates that will cause coking of the beds.
    Conversion to ethylene glycol injection, instead of methanol for hydrate control, will increase sieve life and add at least 10% to sieve capacity by removing the vapor phase methanol from the system.

    5.23.6 Useful Life
    The most common reasons for replacing a bed are loss of adsorbent capacity and unacceptable pressure drop, which usually occur simultaneously. Values for the loss of capacity with time vary considerably, but common values used for molecular sieves in dehydration service are a 35% capacity loss over a 3 to 5 year period or a 50% loss in approximately 1,600 cycles. Typically, a rapid loss occurs in the beginning and a gradual loss thereafter. The adsorbent decays primarily because of carbon and sulfur fouling and caking caused by instability in the clay binder. These effects occur during bed regeneration.
    Increased pressure drop is usually caused by breakdown of adsorbent into finer particles and by caking at the top of the bed because of refluxing. Attrition can occur when the pressure is increased or decreased (more than the allowable pressure-changing rate) after or before regeneration.
    Monitoring the pressure drop is important, as it provides a good diagnostic to bed health.
    In general, adsorbent life ranges from one to four years in normal service. Longer life is possible if feed gas is kept clean. Effectiveness of regeneration plays a major role in retarding the decline of a desiccants adsorptive capacity and prolonging its useful life. If all the water is not removed from the desiccant during each regeneration, its usefulness will sharply decrease.

    5.23.7 Effect of Insufficient Reactivation
    Insufficient reactivation can occur if the regeneration gas temperature or velocity is too low.
    A desiccant manufacturer will generally recommend the optimum regeneration temperature and velocity for the product. Velocity should be high enough to remove the water and other contaminants quickly, thus minimizing the amount or residual water and protect the desiccant.

    5.23.8 Effect of High Reactivation Temperature
    Higher reactivation temperatures remove volatile contaminants before they form coke on the desiccant, maximizes desiccant capacity and ensures minimum effluent moisture content.
    Final effluent hot gas temperature should be held one or two hours to achieve effective desiccant

    5.24 Design
    5.24.1 Pressure Drop & Minimum Diameter
    The first step is to determine the bed diameter, which depends on the superficial velocity. Too large a diameter will require a high regeneration gas rate to prevent channeling. Too small a diameter will cause too high a pressure drop and damage the sieve. The pressure drop is determined by a modified Ergun equation, which relates pressure drop to superficial velocity as follows:
    ΔP / L = B ս V + C ρ V2 Eq 5-17

    ΔP = pressure drop, psi
    L = length of packed bed, ft
    B & C = constants from table 5-10
    = viscosity, cp
    V = superficial vapor velocity, ft/min
    ρ = density, lb/ft3
    Particle Type B C
    1/8" bead (4x8 mesh) 0.0560 0.0000889
    1/8" extrudate 0.0722 0.000124
    1/16" bead (8x12 mesh) 0.152 0.000136
    1/16" extrudate 0.238 0.000210

    Table. 5-10. Constants for Equation 5-6.

    Fig. 5-55 was derived from Eq 5-17 by assuming a gas composition and temperature and setting the maximum allowable ΔP/L equal to 0.33 psi/ft. The design pressure drop through the bed should be about 5 psi. A design pressure drop higher than 8 psi is not recommended as the desiccant is fragile and can be crushed by the total bed weight and pressure drop forces.

    Fig. 5-55. Allowable Velocity for Mole Sieve Dehydrator
    Remember to check the pressure drop after the bed height has been determined. Once the allowable superficial velocity is estimated, calculate the bed minimum diameter
    (i.e., Dminimum), and select the nearest standard diameter (i.e. Dselected):

    Dminimum = (4q /π Vmax)0.5 Eq. 5-18

    D = diameter, ft
    q = actual gas flow rate, ft3/min

    q = m/60ρ Eq. 5-19

    m = mass flow rate, lb/hr
    ρ= density, lb/ft3

    Obtain the corresponding superficial velocity, Vadjusted as follows:
    Vadjusted = Vmax (Dminimum/ Dselected)2 Eq. 5-20
    An alternative and more exact method to calculate the maximum superficial velocity can be determined by Eq 5-21, which was derived from Eq.5-17.
    Vmax = [(ΔP/L)max/(Cρ)]0.5 [(B/C)(ս/ρ)/2] Eq 5-21

    The value of (ΔP/L)max in these equations depends on the sieve type, size, and shape, but a typical value for design is 0.33 psi/ft.

    5.24.2 Mass desiccant Required & Bed Length
    Choose an adsorption period and calculate the mass of desiccant required. Eight to twelve hour adsorption periods are common. Periods of greater than 12 hours may be justified especially if the feed gas is not water saturated.
    Long adsorption periods mean less regenerations and longer sieve life, but larger beds and additional capital investment.
    During the adsorption period, the bed can be thought of as operating with three zones.
    Saturation or equilibrium zone
    The middle or mass transfer zone (MTZ) is where the water content of the gas is reduced from its inlet concentration to < 1 ppm. (lb/MMscf ~ ppmv / 21.4)
    The bottom zone is unused desiccant and is often called the active zone. If the bed operates too long in adsorption, the mass transfer zone begins to move out the bottom of the bed causing a breakthrough.
    In the saturation zone, molecular sieve is expected to hold approximately 13 pounds of water per 100 pounds of sieve. New sieve will have an equilibrium capacity near 20%; 13% represents the approximate capacity of a 3-5 year old sieve.
    This capacity needs to be adjusted when the gas is not water saturated or the temperature is above 75F. Figures. 5-56 and 5.57 are used to find the correction factors for molecular sieve.

    Alternatively, both parameters may be calculated using the following corellations:
    CSS = 0.636 + 0.0826 ln(Sat) Eq. 5-22
    CT = 1.20 0.0026 t(F) Eq. 5-23
    CT = 1.11 0.0047 t(C) Eq. 5-24

    where CSS and CT are correction factors for subsaturation and temperature, respectively.
    Sat is the percent of saturation.
    To determine the mass of desiccant required in the saturation zone, calculate the amount of water to be removed during the cycle and divide by the effective capacity.

    SS = Wr / [(0.13)(CSS)(CT)] Eq 5-25
    LS = 4 SS /[π (D2) (bulk density) Eq 5-26
    LS = length of packed bed saturation zone, ft
    SS = amount molecular sieve required in saturation zone, lb
    Wr = water removed per cycle, lb
    CSS = saturation correction factor for sieve (Fig.5-56)
    CT = temperature correction factor (Fig.5-57)
    LS = length of packed bed saturation zone, ft
    D = diameter, ft

    Molecular sieve bulk density is 42 to 46 lb/ft3 for spherical particles and 40 to 44 lb/ft3 for extruded cylinders.
    Even though the MTZ will contain some water (approximately 50% of the equilibrium capacity), the saturation zone is estimated assuming it will contain all the water to be removed.
    The length of the mass transfer zone can be estimated as follows:

    LMTZ = (Vadjusted/35)0.3 (ZL) Eq 5-27

    LMTZ = length of packed bed mass transfer zone, ft
    ZL = 1.70 ft for 1/8 inch sieve
    = 0.85 ft for 1/16 inch sieve

    Fig. 5-56. Mole Sieve Capacity Correction for gas percent saturation

    Fig. 5-57. Mole Sieve Capacity Correction for Temperature

    The total bed height is the summation of the saturation zone and the mass transfer zone heights. It should be no less than the vessel inside diameter, or 6 feet, whichever is greater.
    Now the total bed pressure drop is checked. The ΔP/L for the selected diameter, Dselected, is adjusted using Eq 5-17 or the following approximation:
    (ΔP / L)adjusted ≅ (0.33 psi/ft) (Vadjusted/Vmax)2 Eq 5-28

    The result is multiplied times the total bed height (LS +LMTZ) to get the total design pressure drop, which should be 5- 8 psi. This is important, because the operating pressure drop can increase to as much as double the design value over three years. Too high a pressure drop plus the bed weight can crush the sieve. If the design pressure drop exceeds 8 psi, the bed diameter should be increased and the sieve amount and vessel dimensions recalculated.
    To estimate the total cylindrical length of a tower, add 3 feet to the bed height, which provides the space for an inlet distributor and for bed support and hold-down balls under and on top of the sieve bed.

    5.24.3 Regeneration Calculations
    The first step is to calculate the total heat required to desorb the water and heat the desiccant and vessel. A 10% heat loss is assumed.
    Qw = (1800 Btu/lb) (lbs of water on bed) Eq 5-29
    Qsi = (lbs of sieve)(0.24 Btu/lb 0F) (Trg Ti) Eq 5-30
    Qst = (lbs of steel)(0.12 Btu/lb 0F) (Trg Ti) Eq 5-31
    Qhl =heat loss = (Qw +Qsi+Qst) (0.1) Eq 5-32

    Qw = desorption of water heat duty, Btu
    Qsi = duty required to heat mole sieve to regeneration temperature, Btu
    Qst = duty required to heat vessel and piping to regeneration temperature, Btu
    Qhl = regeneration heat loss duty, Btu
    Trg = regeneration temperature, 0F
    Ti = initial, temperature, 0F

    The temperature, Trg, is the temperature to which the bed and vessel must be heated based on the vessel being externally insulated (i.e., no internal insulation which is usually the case). This is about 50F below the temperature of the hot regeneration gas entering the tower.
    The weight of the vessel steel is estimated from equations 5-33 and 5-34.
    The design pressure, Pdesign, is usually set at 110% of the maximum operating pressure. The value of 0.125 in Eq 5-34 is the corrosion allowance in inches. The term 0.75Dbed is to account for the weight of the tower heads. The value of "3" provides the space for the inlet distributor and support and hold-down balls.
    t(inches) = (12DbedPdesign) / (37,600 1.2Pdesign) Eq 5-33

    Weight of steel (lb) = 155 (t + 0.125) (LS + LMTZ + 0.75Dbed + 3)Dbed Eq 5-34

    For determination of the regeneration gas rate, calculate the total regeneration load from Eq. 5-35
    Qtr = (2.5)(Qw + Qsi + Qst + Qhl) Eq. 5-35
    Qtr = total regeneration heat duty, Btu
    The 2.5 factor corrects for the change in temperature difference (in out) across the bed with time during the regeneration period. It assumes that 40% of the heat in the gas transfers to the bed, vessel steel, and heat loss to atmosphere; and the balance leaves with the hot gas.

    The regeneration-gas flow rate is calculated from Eq 5-36 below. The heat capacity, Cp, is calculated with Eq 5-37, with the enthalpies obtained from the enthalpy vs. temperature plots for various pressures in (GPSA- Section 24 or refer to definition of (Cp) in eq. 5-37). The temperature, Thot, is 50F above the temperature, Trg, to which the bed must be heated. The temperature, Tb, is the bed temperature at the beginning of regeneration, which is the same as the dehydration-plant feed temperature.
    The heating time is usually 50% to 60% of the total regeneration time which must include a cooling period. Figure 5-59 shows a typical temperature profile for a regeneration period (heating and cooling). For 8 hour adsorption periods, the regeneration normally consists of 4 1/2 hours of heating, 3 hours of cooling and 1/2 hour for standby and switching. For longer periods the heating time can be lengthened as long as a minimum pressure drop of 0.01 psi/ft is maintained to ensure even flow distribution across the bed.
    mrg = Qtr/[Cp(Thot -Tb)(heating time)] Eq 5-36

    Cp (Btu/lb/F) = (Hhot - Hi)/(Thot -Tb) Eq 5-37
    mrg = regeneration mass flow rate, lb/hr
    Thot = Hot gas temperature, 0F
    Tb = bed starting temperature, 0F
    Cp = gas heat capacity, Btu/(lb . F) = (Hhot - Hi)/(Thot -Tb) (Extract the value from fig. 5-60 ,which include values at 600 psia, or (Use Cp = 0.66 as an approximate value for natural gas in regeneration operation ranges) Hhot = enthalpy, BTU/lb, and Hi = enthalpy, BTU/lb. (Other values can be extracted from curves in GPSA, chapter 24 Total Enthalpy of Paraffin Hydrocarbon Vapor).
    The superficial velocity of the regeneration gas is calculated from Eq 5-38 for which (q) is calculated from Eq 5-19
    V = 4q / (π D2) Eq 5-38
    V = Velocity, ft/sec
    D = diameter, ft
    q = actual gas flow rate, ft3/min
    The calculated superficial velocity can not be less than the value that corresponds with a minimum bed pressure drop of 0.01 psi/ft. This can be determined from Fig. 5-58, which was derived from Eq 5-17 by assuming a gas composition and temperature and setting ΔP/L equal to 0.01 psi / ft.
    If the calculated velocity is less than this, the regeneration gas rate, (mrg) must be increased by multiplying it by the ratio Vmin / V, and the period of regeneration should be decreased by multiplying it times the ratio V / Vmin. A more exact method for calculating the minimum superficial velocity is to use Eq 5-21, but to consider it in terms of (ΔP/L)min and Vmin instead of (ΔP/L)max and Vmax.

    Fig. 5-58. Minimum Regeneration Velocity for Mole Sieve Dehydrator

    Fig. 5-59. Inlet and Outlet Temperatures During Typical Solid Desiccant Bed Regeneration Period
    5.24.4 Design Example
    100 MMscfd of natural gas with a molecular weight of 18 is water saturated at 600 psia and 100F and must be dried to 150F dew point.
    Determine the water content of the gas, and the amount of water that must be removed; and do a preliminary design of a molecular-sieve dehydration system consisting of two towers (one dehydration and the other regeneration in a cycle basis). Use 4A molecular sieve of 1/8" beads (i.e., 4x8 mesh). Compressibility factor z = 0.93.
    The regeneration gas is part of the plants residue gas, which is at 600 psia and 100F and has a molecular weight of 17. The bed must be heated to 500F for regeneration. Compressibility factor = 0.95.

    Solution Steps
    1. Determine the bed diameter and the corresponding ΔP/L and V. First determine the maximum superficial velocity from Eq 5-10. Let the maximum ΔP/L be 0.33 psi/ft.
    Vmax = [(ΔP/L)max/(Cρ)]0.5 [(B/C)(ս/ρ)/2] Eq 5-21

    Fro chapter 1, Eq. 1-9
    ρg= 0.093 ((MW)P)/TZ lb/ft3 (Eq. 1-19)
    ρg = 0.093 x 18 x 600 / (560 x 0.93)
    ρ = 1.93 lb/ft3
    ρ = density of gas (ρ)= (18 mole weight) (600 psia) / [10.73 (560 R)(0.93 z)] = 1.93 lb/ft3
    = 0.015 centipoise (Fig. 1-11)
    C = 0.000089 from table 5-10.
    Vmax = {(0.33 psi/ft) / (0.0000889)(1.93) lb/ft3)}0.5 [(0.056 / 0.0000889) (0.015 centipoise/1.93 lb/ft3)/2]
    = 41.4 ft/min
    Mass flow rate (m) = (100 X 106 scf/day) (18 lb/lb mole) / [(24 hr/day)(379.5 scf/lb mole)]
    m = 198,000 lb/hr of wet gas
    From Eq. 5-19
    q = (198,000 lb/hr)/((60 min/hr)(1.93 lb/ft3)) = 1710 ft3/min of wet gas
    From eq. 5-18
    Dminimum = [(4(1710 ft3/min))/(3.14 X 41.4 ft/min)]0.5 = 7.25 ft

    Round off upward to 7.5 ft diameter, for which V and ΔP/L are adjusted as follows (eq. 5-20):
    Vadjusted = (41.4)(7.25/7.5)2 = 38.7 ft/min
    (ΔP/L)adjusted = 0.33(38.7/41.4)2 = 0.29 psi/ft

    2. Estimate the amount of water to be removed from the feed per cycle for each bed.
    Base this on a 24-hour cycle consisting of 12 hours adsorbing and 12 hours regenerating (heating, cooling, standby, and valve switching).
    From chapter 4, fig. 4-8, the water content at 600 psia and 100F is 88 pounds/MMscf. The water content at a dew point of 150F is essentially zero, so the water removed is the following:
    w = (88-0 lb/MMscf)(100 MMscf/day) / (24 hr/day)
    = 367 lb/hr of water removed
    Wr = (367 lb/hr) (12 hr) = 4404 lb water removed per 12-hour drying period or 24-hour cycle per bed.

    3. Determine the amount of sieve required and the bed height based on a sieve bulk density of 45 lb/ft3 (table 5-7). Since the feed gas is water saturated, the relative humidity is 100%, so CSS is 1.0, and CT is 0.93 at 100F from Figures. 5-56 & 5-57.
    Applying the equations 5-25:
    SS = Wr / [(0.13)(CSS)(CT)] Eq 5-25
    SS = (4404) / ((0.13)(1.0)(0.93)) = 36,427 lb of sieve for each bed.
    From eq. 5-26.
    LS = 4 SS /[π (D2) (bulk density) Eq 5-26
    LS = (4)(36,427) /((3.1416) (7.5)2 (45)) = 18.3 ft bed height
    from eq-5-27
    LMTZ = (Vadjusted/35)0.3 (ZL) Eq 5-27
    LMTZ = (38.7/35)0.3 (1.7) = 1.8 ft for mass-transfer zone
    LS + LMTZ = 18.3 + 1.8 = 20.1 ft of sieve for each bed

    The total sieve = (20.1/18.3)(36,427) = 40,010 lb for each bed

    4. Check the bed design and pressure drop which is the ΔP/L calculated in Step 1 times the total bed height calculated in Step 3:
    (0.29 psi/ft)(20.1 ft) = 5.8 psi which meets the criterion of not exceeding 8 psi.

    5. Calculate the total heat required to desorb the water based on heating the bed and vessel to 500F. First calculate the weight of steel from Eq 5-33 and 5-34. Let the design pressure, Pdesign, be 110% of the operating pressure: Pdesign = (600)(1.1) = 660 psia.
    t(inches) = (12DbedPdesign) / (37,600 1.2Pdesign) Eq 5-33
    Weight of steel (lb) = 155 (t + 0.125) (LS + LMTZ + 0.75Dbed + 3)Dbed Eq 5-34

    t = (12)(7.5)(660) / (37,600 (1.2)(660)) = 1.614 inches
    Weight of steel = (155) (1.614 + 0.125) (18.3 + 1.8)+ (0.75) (7.5) + 3) (7.5) = 58,070 pounds
    From the following equations:
    Qw = (1800 Btu/lb) (lbs of water on bed) Eq 5-29
    Qsi = (lbs of sieve)(0.24 Btu/lb 0F) (Trg Ti) Eq 5-30
    Qst = (lbs of steel)(0.12 Btu/lb 0F) (Trg Ti) Eq 5-31
    Qhl =heat loss = (Qw +Qsi+Qst) (0.1) Eq 5-32
    Qtr = (2.5)(Qw + Qsi + Qst + Qhl) Eq. 5-35

    Qw = (1800 Btu/lb (4404 lb water) = 7,927,000 Btu
    Qsi = (40,010 lb) (0.24 Btu/lb/F) (500F 100F) = 3,841,000 Btu
    Qst = (58,070 lb) (0.12 Btu/lb/F) (500F 100F) = 2,787,000 Btu
    Qhl = (2,787,000 + 7,927,000 + 3,841,000) (0.10) = 1,455,000 Btu
    Qtr = (2.5) (2,787,000 + 7,927,000 + 3,841,000 + 1,455,000) = 40,025,000 Btu

    6. Calculate the flow rate of regeneration gas using Eq 5-36. Let the heating time be 60% of the total regeneration period, and calculate the gas heat capacity, use Cp, = 0.66 (Btu/lb/F), or (calculate it from Eq 5-37 using enthalpy curves in GPSA, chapter 24 Total Enthalpy of Paraffin Hydrocarbon Vapor):
    mrg = Qtr/[Cp(Thot -Tb)(heating time)] Eq 5-36
    Cp (Btu/lb/F) = (Hhot - Hi)/(Thot -Tb) Eq 5-37

    (60%) (12 hr) = 7.2 hours heating
    Cp (@ 600 psia from fig 5-60) = ((545 250) (Btu/lb))/((550 100) (F)) = 0.66 Btu/lb/F

    mrg = (40,025,000 Btu)/((0.66 Btu/lb/F) (550 100) (F) (7.2 hr)) = 18,717 lb/hr (Eq 5-22)

    7. Check that the ΔP/L ≥ 0.01 psi/ft at 550F.
    ρg = 0.093 x 17 x 600 / (1110 x 0.95)
    ρ = 0.9 lb/ft3

    From Eq. 5-19 q = m/60ρ
    = 18,717 / (60 x 0.9) = 346.6 ft3/min of hot regeneration gas

    Rearranging Eq 5-38:
    V = 4q / (π D2) Eq 5-38
    V = 4q/πD2 = ((4)(346.6)/((3.414)(7.5)2) = 7.21 ft/min
    μ = 0.023 cP (Fig. 1-11)
    From ΔP / L = B ս V + C ρ V2 Eq 5-17
    ΔP/L = (0.056) (0.023) (7.21) + (0.0000889) (0.9) (7.21)2 = 0.013 psi/ft (Eq 5-6)
    This is safely above the minimum value of 0.01 psi/ft needed to prevent channeling.

    8. The design results are summarized as follows:
    Number of vessels: two
    Vessel design pressure and temperature: 660 psig and 600F
    Vessel dimensions: 90 inches (7.5 feet) ID by 23.1 feet tan to tan
    Weight of molecular sieve: 2x40,010 lb
    Regeneration gas rate: 18,717 lb/hr (10.026 MMscfd)
    Regeneration gas temperature: 550F
    Cycle time: 24 hours, 12 hours adsorption, 12 hours regeneration
    Fig.5-60. Total Enthalpy of Paraffin Hydrocarbon Vapor @ 600 psia.

    5.25 Nonregenerable Dehydrator
    In some situations, such as remote gas wells, use of a consumable salt desiccant, such as CaCl2, may be economically feasible.

    5.25.1 Calcium Chloride Dehydrator Unit
    Calcium chloride (CaCI2) dehydrator consists of three sections (Figure 5-61):
    Inlet gas scrubber
    Brine tray
    Solid brine particles

    5.25.2 Principles of Operation
    Solid desiccant is placed in the top of the unit. Water-wet gas contacting the solid CaCI2 gives up part of its water to form liquid brine to drip down and fill the trays.
    Solid anhydrous CaCl2 combines with water to form various CaCl2 hydrates (CaCl2 .XH2O). As water absorption continues, CaCl2 is converted to successively higher states of hydration eventually forming a CaCl2 brine solution.
    Inlet gas coming up through the specially designed nozzles on the trays contact the brine efficiently. The wettest gas contacts the most dilute brine (about 1.2 specific gravity).
    Approximate 2.5 Lb H2O/lb CaCI2 is removed in the trays. Brine gravity on the top tray is about 1.4. Another 1 lb H2O/lb CaCI2 is removed in the solid bed section.
    Outlet water contents of 1 lb/MMscf have been achieved with CaCl2 dehydrators. Typical CaCl2
    Maximum dew point depression of 60 0 to 70 0F occurs in this section. Typically used in remote, small gas fields without heat or fuel. capacity is 0.3 lb CaCl2 per lb H20. Superficial bed velocities are 20-30 ft/min and length to diameter ratio for the bed should be at least 3 to 4:1.
    CaCl2 dehydrators may offer a viable alternative to glycol units on low rate, remote dry gas wells. The CaCl2 must be changed out periodically. In low capacity high rate units this may be as often as every 2-3 weeks. Brine disposal raises environmental issues. In addition, under certain conditions the CaCl2 pellets can bond together to form a solid bridge in the fixed bed portion of the tower. This results in gas channeling and poor unit performance.

    Advantages Disadvantages
    No moving parts
    No heat required
    Does not react with H2S or CO2
    Can dehydrate hydrocarbon liquids Batch process
    Emulsifies with oil
    Limited dew point depression

    Table. 5-11. Advantages and disadvantage of Calcium Chloride Dehydrator Unit

    Operating Problems
    Bridging and channeling is a problem.
    Brine can crystallize at 85 0F, thus during low flow periods can plug vessel outlet or trays.
    Brine carry-over can cause severe corrosion problems.

    Fig. 5-61. CaCl2 dehydrator.

    Design Considerations
    Figure 5-62 illustrates the water content of natural gas dried by solid calcium chloride bed units.

    Fig. 5-62 Water content on natural gas driven by CaCl2 unit (Left: freshly recharged; right: just prior to recharging).
    5.26 Dehydration by Refrigeration
    The dehydration of natural gas can also be achieved by refrigeration and/or cryogenic processing down to 150F in the presence of methanol hydrate and freeze protection. The condensed water and methanol streams decanted in the cold process can be regenerated by conventional distillation or by a patented process called IFPEX-1.
    In the latter process illustrated in schematic form in Figure 5-63 a slip stream of water saturated feed gas strips essentially all the methanol in the cold decanted methanol water stream originating in the cold process at feed gas conditions to recirculate the methanol to the cold process. The water stream leaving the stripper contains generally less than 100 ppm wt of methanol. No heat is required for the process and no atmospheric venting takes place.
    The process has several major advantages:
    It can obtain dew points in the −100 to −150F (−70 to 100C) range.
    It requires no heat input other than to the methanol regenerator.
    It requires no venting of hydrocarbon-containing vapors.
    However, it requires external refrigeration to cool the gas, and minimal methanol losses occur in the stripper.

    Fig. 5-63. Example IFPEX-1 . Dehydration Process Flow Diagram

    5.27 Dehydration by Membrane Permeation
    Membranes can be used to separate gas stream components in natural gas such as water, CO2 and hydrocarbons according to their permeabilities. Each gas component entering the separator has a characteristic permeation rate that is a function of its ability to dissolve in and diffuse through the membrane.
    The driving force for separation of a gas component in a mixture is the difference between its partial pressure across the membrane. As pressurized feed gas flows into the metal shell of the separator, the fast gas component, such as water and CO2, permeate through the membrane. This permeate is collected at a reduced pressure, while the non-permeate stream, i.e., the dry natural gas, leaves the separator at a slightly lower pressure than the feed.
    The amount of methane and other natural gas components in the permeate stream is dependent on pressure drop and the surface area of the membranes. However, 510% of the feed stream is a realistic figure. Dehydration by membrane permeation is therefore normally only considered for plants that can make use of low pressure natural gas fuel.

    Membranes are characteristic by lightweight, large turndown ratio, and low maintenance, that make them competitive with glycol units in some situations.
    Feed pretreatment is a critical component of a membrane process. The inlet gas must be free of solids and droplets larger than 3 microns.
    Inlet gas temperature should be at least 20F (10C) above the dew point of water to avoid condensation in the membrane.
    Units operate at pressures up to 700 to 1,000 psig (50 70 barg) with feed gases containing 500 to 2,000 ppmv of water (lb/MMscf ~ ppmv / 21.4). They produce a product gas stream of 20 to 100 ppmv and 700 to 990 psig (48 to 68 barg). The low-pressure (7 to 60 psig [0.5 to 4 barg]) permeate gas volume is about 3 to 5% of the feed gas volume.
    This gas must be recompressed or used in a low-pressure system such as fuel gas.
    Smith (2004) suggests that membranes used for natural gas dehydration are economically viable only when dehydration is combined with acid-gas removal.
    On the basis of commercial units installed and several studies; (Bikin et al., 2003), membranes are economically attractive for dehydration of gas when flow rates are less than 10 MMscfd (0.3 MMSm3/d). Binci et al. claim that membrane units are competitive with TEG dehydrators on offshore platforms at flows below 56 MMscfd (1.6 MMSm3/d). Certainly, the reliability and simplicity of membranes make them attractive for offshore and remote-site applications, provided the low-pressure permeate gas is used effectively. An added benefit compared with TEG units is the absence of BTEX emissions with membranes.
    5.28 Other Processes
    The first process is the Twister technology, which is discussed in Chapter 6. It has been considered attractive in offshore applications for dehydration because of its simplicity (no moving parts) along with its small size and weight.
    Brouwer et al. (2004) discuss the successful implementation on an offshore platform. Some offshore field pressures are greater than 2,000 psi (140 bar), so recompression is not needed with the unit where overall pressure drop is 20 to 30%.
    The second process is the vortex tube technology, which also is discussed in Chapter 6. It also has no moving parts. According to vendor information, it is used in Europe in conjunction with TEG addition to remove water from gas stored underground. We found no examples of its use in gas plants.
    5.29 Comparison of Dehydration Processes
    A number of factors should be considered in the evaluation of a dehydration process or combination of processes. If the gas must be dried for cryogenic liquids recovery, molecular sieve is the only long-term, proven technology available. It has the added advantage that it can remove CO2 at the same time. If CO2 is being simultaneously removed, because water displaces CO2, the bed must be switched before the CO2 breaks through, which is before any water breakthrough.

    Enhanced TEG regeneration systems may begin to compete with molecular sieve. Skiff et al. (2003) claim to have obtained less than 0.1 ppmv water by use of TEG with a modified regeneration system that uses about 70% of the energy required for molecular sieves.

    High inlet water-vapor concentrations make molecular sieve dehydration expensive because of the energy consumption in regeneration. Two approaches are used to reduce the amount of water going to the molecular sieve bed. First, another dehydration process, (e.g., glycol dehydration) is put in front of the molecular sieve bed. The second option is to have combined beds with silica gel or activated alumina in front of the molecular sieve. The bulk of the water is removed with the first adsorbent, and the molecular sieve removes the remaining water. This configuration reduces the overall energy required for regeneration.

    If dehydration is required only to avoid free-water formation or hydrate formation or to meet the pipeline specification of 4 to 7 lb/MMscf (60 to 110 mg/Sm3), any of the above-mentioned processes may be viable. Traditionally, glycol dehydration has been the process of choice.

    System constraints dictate which technology is the best to use. Smith (2004) provides an overview of natural gas dehydration technology, with an emphasis on glycol dehydration.
    When considering susceptibility to inlet feed contamination, one should keep in mind that replacing a solvent is much easier and cheaper than changing out an adsorbent bed. However, prevention of contamination by use of properly designed inlet scrubbers and coalescing filters, if required, is the best solution.
    In a conventional gas plant, where inlet fluctuations are handled in inlet receiving, feed contamination is generally limited to possible carryover from the sweetening unit. However, in field dehydration the possibility exists of produced water, solids, oil, and well-treating chemicals entering the dehydrator.

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    Gas Sweetening and Sulfur Recovery - Chapter 6 - Part 1 A

    Chapter 6 216
    Gas Sweetening 216
    6.1 Introduction 216
    6.1.1 Definition 216
    6.1.2 Safety Precautions 216
    6.1.3 Purification Levels 217
    6.1.4 Acid Gas Disposal 217
    6.1.5 Removal of contaminants and pretreatment 218
    6.2 Gas Treating Process Options 219
    6.3 Chemical Solvent Processes 221
    6.3.1 Amine (Aqueous Alkanolamine) Processes 223
    6.3.2 Alkaline Salt Process (Hot Carbonate) 234
    6.3.3 Specialty Batch Chemical Solvents 237
    6.4 Physical Solvent Processes 238
    6.4.1 Selexol 239
    6.4.2 Fluor Solvent 242
    6.4.3 Rectisol Process 243
    6.4.5 Purisol 243
    6.4.5 Catasol 243
    6.4.6 Morphysorb 243
    6.5 Hybrid Solvent Processes 244
    6.5.1 Sulfinol Process 244
    6.5.2 Selefining Process 244
    6.6 General Considerations for Solvent Process 245
    6.6.1 Solution Filtration 245
    6.6.2 Flash Tank 245
    6.6.3 Corrosion 246
    6.6.4 Foaming 247
    6.6.5 Materials 248
    6.7 Solid Bed Processes 248
    6.7.1 General Process Description 248
    6.7.2 Iron Sponge Process 248
    6.7.3 SulfaTreat 250
    6.7.4 Zinc Oxide Process 251
    6.7.5 Chemsweet 251
    6.7.6 PuraSpec 251
    6.7.8 Molecular Sieve Process 251
    6.7.9 Oxorbon 253
    6.8 Direct Conversion Processes (Liquid Redox) 254
    6.8.1 Stretford Process 255
    6.8.2 Lo-Cat process 255
    6.8.3 Sulferox process 257
    6.8.4 IFP Process 258
    6.9 Distillation Process 258
    6.10 Sulfur Recovery (The Claus Process) 259
    6.10.1 Claus Process Considerations 261
    6.10.2 Process Variations 262
    6.10.3 Combustion Operation 263
    6.10.4 Claus Unit Tail Gas Handling 263
    6.11 Gas Permeation Process (Membranes) 266
    6.11.1 Membrane Fundamentals 266
    6.11.2 Membrane Selection Parameters 266
    6.11.3 Membrane Structure Types 267
    6.11.4 Carbon Dioxide Removal from Natural Gas 268
    6.11.5 Membrane Elements 268
    6.11.6 Membrane Design Considerations 271
    6.11.7 Operating Considerations 274
    6.11.8 Feed Gas Pretreatment 277
    6.11.9 Membrane Advantages & Disadvantages 280
    6.11.10 Hybrid Configurations 282
    6.12 Biological Processes 285
    6.13 Process Selection 285
    6.13.1 Inlet Gas Stream Analysis 285
    6.13.2 General Considerations 285
    6.13.3 Removal of H2S 286
    6.13.4 Removal of H2S and CO2 286
    6.13.5 Process Selection charts 287
    6.14 Safety & Environmental Considerations 289
    6.15 Design Procedure 290
    6.15.1 Iron Sponge 290
    Example 6-1 291
    6.15.2 The Amine System 292
    Method 1 292
    Example 6-2 294
    Method 2 294
    Example 6-3 296

    Chapter 6

    Gas Sweetening

    6.1 Introduction
    Gas treating involves reduction of the acid gases carbon dioxide (CO2) and hydrogen sulfide (H2S), along with other sulfur species, to sufficiently low levels to meet contractual specifications or permit additional processing in the plant without corrosion and plugging problems. This chapter focuses on acid gases because they are the most prevalent. When applicable, discussion is given to other sulfur species.
    There are many methods that may be employed to remove acidic components (primarily H2S and CO2) and other impurities from hydrocarbon streams. The available methods may be broadly categorized as those depending on chemical reaction, absorption, adsorption or permeation. Processes employing each of these techniques will be described.

    6.1.1 Definition
    Acid Gases
    H2S combined with water forms sulfuric acid. CO2 combined with water forms carbonic acid.
    Both are undesirable because they cause corrosion and reduce heating value and sales value. H2S is poisonous and may be lethal.
    Sour Gas:
    Sour gas is defined as natural gas with H2S and other sulfur compounds.
    Sweet Gas:
    Sweet gas is defined as natural gas without H2S and other sulfur compounds.
    Acid gases Partial Pressure:
    Partial pressure is used as an indicator if treatment is required. Partial pressure is defined as
    PP = (total pressure of system) x ( mol% of gas)
    where CO2 is present with water, a partial pressure >30 psia, would indicate CO2 corrosion might be expected. Below 15 psia, would indicate CO2 corrosion would not normally be a problem although inhibition may be required.
    Factors that influence CO2 corrosion are those directly related to solubility, that is, temperature, pressure, and composition of the water. Increased pressure increases solubility and increased temperature decreases solubility.
    H2S may cause sulfide stress *****ing due to hydrogen embrittlement in certain metals. H2S partial pressure >0.05 psia, necessitates treating.

    6.1.2 Safety Precautions
    Hydrogen sulfide is a highly toxic gas. At concentrations as low as 10 ppmv irritation of the eyes, nose, and throat is possible. The human nose can detect hydrogen sulfide in concentrations as low as 0.02 ppmv. However, the human sense of smell cannot be relied on to detect hazardous concentrations of hydrogen sulfide. Higher concentrations and extended exposure to hydrogen sulfide will desensitize the sense of smell. The concentrations required for different reactions by the human body are:
    1. Threshold limit value (TLV) for prolonged exposure: 10 ppmv
    2. Slight symptoms after several hours exposure: 10-100 ppmv
    3. Maximum concentration that can be inhaled for one hour without serious effects such as significant eye and respiratory irritation: 200-300 ppmv
    4. Dangerous after exposure of 30 minutes to one hour: 500-700 ppmv
    5. Fatal in less than 30 minutes: 700-900 ppmv and above.
    6. Death in minutes: greater than 1000 ppmv

    Hydrogen sulfide is highly flammable and will combust in air at concentrations from 4.3 to 46.0 volume percent. Hydrogen sulfide vapors are heavier than air and may migrate considerable distances to a source of ignition.
    When H2S concentrations are well above the ppmv level, other sulfur species can be present. These compounds include carbon disulfide (CS2), mercaptans (RSH), and sulfides (RSR), in addition to elemental sulfur. If CO2 is present as well, the gas may contain trace amounts of carbonyl sulfide (COS). The major source of COS typically is formation during regeneration of molecular-sieve beds used in dehydration (see Chapter 5).

    Gaseous carbon dioxide is a naturally occurring gas that is 50% heavier than air and is colorless and odorless. It is also a principal by-product of combustion. CO2 is inactive and therefore non-flammable. CO2 will displace oxygen and can create an oxygen-deficient atmosphere resulting in suffocation. The principal hazard of CO2 is exposure to elevated concentrations.
    The atmospheric concentration immediately hazardous to life is 10%. Because CO2 is heavier than air, its hazard potential is increased, especially when entering tanks and vessels.
    "A common but erroneous belief is that CO2 simply acts as an asphyxiant by lowering the oxygen level below the 16 percent minimum necessary to sustain life (at sea level). Although this is frequently the case in most serious accidents, CO2 begins to have a noticeable effect on normal body functions at about two to three percent. The concentration of carbon dioxide in the blood affects the rate of breathing, a measurable increase resulting from a level of one percent in the inspired air."
    6.1.3 Purification Levels
    Although many natural gases are free of objectionable amounts of H2S and CO2, substantial quantities of these impurities are found in both gas reserves and production. In a survey of U.S. gas resources, Meyer (2000) defined subquality gas as that containing CO2 ≥ 2%, N2 ≥ 4%, or H2S ≥ 4 ppmv.
    The inlet conditions at a gas processing plant are generally temperatures near ambient and pressures in the range of 300 to 1,000 psi (20 to 70 bar), so the partial pressures of the entering acid gases can be quite high. If the gas is to be purified to a level suitable for transportation in a pipeline and used as a residential or industrial fuel, then the H2S concentration must be reduced to 0.25 gr/100 scf (4 ppm, or 6 mg/m3), and the CO2 concentration must be reduced to a maximum of 3 to 4 mol%. However, if the gas is to be processed for NGL recovery or nitrogen rejection in a cryogenic turbo expander process, CO2 may have to be removed to prevent formation of solids. If the gas is being fed to an LNG liquefaction facility, then the maximum CO2 level is about 50 ppmv because of potential solids formation.

    6.1.4 Acid Gas Disposal
    What becomes of the CO2 and H2S after their separation from the natural gas?
    The answer depends to a large extent on the quantity of the acid gases. For CO2, if the quantities are large, it is sometimes used as an injection fluid in EOR (enhanced oil recovery) projects. Several gas plants exist to support CO2 flooding projects. If this option is unavailable, then the gas can be vented, provided it satisfies environmental regulations for impurities.
    In the case of H2S, four disposal options are available:
    1. Incineration and venting, if environmental regulations regarding sulfur dioxide emissions can be satisfied
    2. Reaction with H2S scavengers, such as iron sponge
    3. Conversion to elemental sulfur by use of the Claus or similar process
    4. Disposal by injection into a suitable underground formation.
    The first two options are applicable to trace levels of H2S in the gas, and the last two are required if concentrations are too high to make the first two options feasible.
    6.1.5 Removal of contaminants and pretreatment
    All gas sweetening units should have well-designed pretreatment facilities. Carryover of brine or liquid hydrocarbon (as slugs or aerosol) from upstream production operations can cause problems for gas treating and downstream processing equipment. Also, field facilities are not typically designed to remove troublesome contaminants like gas-phase heavy hydrocarbons.
    These contaminants can likewise cause operational difficulties in the sweetening process.

    If gross liquid carryover from an upstream facility is possible, a slug catcher is recommended. It should be sized not only for steady inlet fluid volumes, but for surge capacity to handle slugs of liquid hydrocarbons, water, and/ or well treatment chemicals.
    If aerosols are a concern, an inlet filter separator is suggested. Selected filter elements can remove entrained droplets down to 0.3 microns in diameter.
    For liquid hydrocarbon treatment, a filter coalescer may be used to remove suspended water or glycol prior to further processing.
    Heavy hydrocarbons (C6+) can be absorbed by solvents, which could lead to foaming in the sweetening unit. It is possible to reduce the heavy hydrocarbon content of the incoming gas through cooling (via Joule-Thomson expansion, propane refrigeration, or turbo-expansion), and subsequent condensation of the heavy components. The condensed liquids are removed, and the gas is warmed above the saturation temperature before going to the sweetening unit.
    An alternative means to remove gaseous heavy hydrocarbons is through adsorption. Either alumina or silica gel beds may be used in parallel such that one bed is regenerated while the other is in service. The beds are regenerated by heating and desorbing the hydrocarbons. The heavy hydrocarbons are recovered from the regeneration gas via condensation.

    Oxygen entry into a hydrocarbon system is often troublesome. If liquid water is present, severe corrosion may occur. If H2S or sulfur is present, corrosion by a different mechanism or sulfur deposition and plugging may occur. Oxygen contamination may be addressed by several different approaches but the first step is to find and correct the source of oxygen entry into the system. This is often the simplest and most cost effective approach. Most oxygen leaks may be traced to compressor suctions or pipe fittings.
    To eliminate oxygen contamination a number of possibilities exist:
    React the oxygen with chemicals.
    Chemicals such as amines, organics or inorganic compounds may be added to remove free oxygen. Oxygen scavengers are available from many chemical suppliers.
    Thermal oxidation reactions.
    Integrated processes such as the DeOxy by Optimized Process Designs can perform a very limited burn to consume the free oxygen.
    Remove other reactants that cause problems with the presence of oxygen.
    By removing offending components such as water or H2S that react with oxygen, the presence of low amounts of oxygen may be tolerated.
    Treat the symptom.
    Corrosion inhibitors, filtration and / or alternate schemes may be utilized to stop or offset any adverse effects of oxygen contamination.
    6.2 Gas Treating Process Options
    Sulfur exists in natural gas as hydrogen sulfide (H2S), and the gas is usually considered sour if the hydrogen sulfide content exceeds 4 ppm of H2S. The process for removing hydrogen sulfide and carbon dioxide from a natural gas stream is referred to as sweetening the gas.
    Numerous processes have been developed for acid gas removal and gas sweetening based on a variety of chemical and physical principles.
    Some of the more important items that must be considered before a process is selected are:
    The type and concentration of impurities and hydrocarbon composition of the sour gas. For example, COS, CS2, and mercaptans can affect the design of both gas and liquid treating facilities. Physical solvents tend to dissolve heavier hydrocarbons, and the presence of these heavier compounds in significant quantities tends to favor the selection of a chemical solvent.
    The temperature and pressure at which the sour gas is available. High partial pressures (50 psi [3.4 bar] or higher) of the acid gases in the feed favor physical solvents, whereas low partial pressures favor the amines.
    The specifications of the outlet gas (low outlet specifications favor the amines).
    The volume of gas to be processed.
    The specifications for the residue gas, the acid gas, and liquid products.
    The selectivity required for the acid gas removal.
    The capital, operating, and royalty costs for the process.
    The environmental constraints, including air pollution regulations and disposal of byproducts considered hazardous chemicals.

    If gas sweetening is required offshore, both size and weight are additional factors that must be considered. Whereas CO2 removal is performed offshore, H2S removal is rarely done unless absolutely necessary because of the problems of handling the rich acid gas stream or elemental sulfur.

    If the gas processing facility is to be used in conjunction with liquids recovery, the requirements for H2S, CO2, and mercaptan removal may be affected. In liquid recovery plants, varying amounts of H2S, CO2, and other sulfur compounds will end up in the liquid product. Failure to remove these components prior to liquids recovery may require liquid product treating in order to meet product specifications. In many instances, liquid treating may be required anyway.
    When sulfur recovery is required, the composition of the acid gas stream feeding the sulfur plant must be considered. With CO2 concentrations greater than 80% in the acid gas, the possibility of selective treating should be considered to raise the H2S concentration to the sulfur recovery unit (SRU). This may involve a multi-stage gas treating system in which the gas exiting the first stage is enriched by passing it through another absorption solvent loop.
    High concentrations of hydrocarbons can cause design and operating problems for the SRU. The effect of these components must be weighed when selecting the gas treating process to be used.
    Decisions in selecting a gas treating process can be simplified by gas composition and operating conditions.
    High partial pressures (50 psi) of acid gases enhance the possibility of using a physical solvent. The presence of significant quantities of heavy hydrocarbons in the feed discourages using physical solvents. Low partial pressures of acid gases and low outlet specifications generally require the use of amines for adequate treating. Process selection is not easy and a number of variables must be weighed prior to making a process selection.
    Figure 6-1 summarizes the more important processes and groups them into the generally accepted categories.

    Fig. 6-1. Acid Gas Removal Processes

    Table 6.1 lists the processes used to separate the acid gas from other natural gas components.
    Tables 6-2 and 6-3, include comparison of different gas sweetening processes.

    Table 6.1 Acid Gas Removal Processes

    Note a: MEA reacts nonreversibly with COS (carbonyl sulfide), and, therefore, should not be used to treat gases with a large concentration of COS.
    Table 6.2 Gases Removed by Various Processes

    6.3 Chemical Solvent Processes
    Chemical reaction processes remove the H2S and/or CO2 from the gas stream by chemical reaction with a material in the solvent solution. The reactions may be reversible or irreversible.
    In reversible reactions the reactive material removes CO2 and/or H2S in the contactor at high partial pressure and/or low temperature. The reaction is reversed by high temperature and/or low pressure in the stripper. In irreversible processes the chemical reaction is not reversed and removal of the H2S and/or CO2 requires continuous makeup of the reacting material.

    Fig. 6-6 shows the process flow for a typical reversible chemical reaction process.
    In solvent absorption, the two major cost factors are the solvent circulation rate, which affects both equipment size and operating costs, and the energy require and disadvantages of chemical and physical solvents. Table 6.4 summarizes some of the advantages and disadvantages of chemical and physical solvents.

    Table 6-3. CO2 and H2S Removal Processes for Gas Streams.

    The most common chemical solvents are:
    Amines (Aqueous Alkanolamine)

    Table 6.4. Comparison of Chemical and Physical Solvents

    6.3.1 Amine (Aqueous Alkanolamine) Processes
    All commonly used amines are alkanolamines, which are amines with OH groups attached to the hydrocarbon groups to reduce their volatility. Figure 6.2 shows the formulas for the common amines used in gas processing.
    Amines are compounds formed from ammonia (NH3) by replacing one or more of the hydrogen atoms with another hydrocarbon group. Replacement of a single hydrogen produces a primary amine, replacement of two hydrogen atoms produces a secondary amine, and replacement of all three of the hydrogen atoms produces a tertiary amine. Primary amines form stronger bases than secondary amines, which form stronger bases than tertiary amines. Amines with stronger base properties are more reactive toward CO2 and H2S gases and form stronger chemical bonds. Sterically hindered amines are compounds in which the reactive center (the nitrogen) is partially shielded by neighboring groups so that larger molecules cannot easily approach and react with the nitrogen. The amines are used in water solutions in concentrations ranging from approximately 10 to 65 wt% amines.(table. 6-5 lists Amines physical properties ).

    Table. 6-6. lists approximate guidelines for a number of alkanolamine processes.
    The amine processes are particularly applicable where acid gas partial pressures are low and/or low levels of acid gas are desired in the residue gas.
    Because the water content of the solution minimizes heavy hydrocarbon absorption, these processes are well suited for gases rich in heavier hydrocarbons. Some amines can be used to selectively remove H2S in the presence of CO2.
    Amines remove H2S and CO2 in a two step process:
    1. The gas dissolves in the liquid (physical absorption).
    2. The dissolved gas, which is a weak acid, reacts with the weakly basic amines.

    Absorption from the gas phase is governed by the partial pressure of the H2S and CO2 in the gas, whereas the reactions in the liquid phase are controlled by the reactivity of the dissolved species. The principal reactions are summarized in the next section.

    Fig. 6.2 Molecular structures of commonly used amines.

    Table 6-5 physical properties of gas sweetening chemicals.
    Table 6-6. Approximate Guidelines for Amine Processes

    Table. 6-5. Physical properties of gas treating chemicals.

    Table 6-6. Approximate Guidelines for Amine Processes
    1. These data alone should not be used for specific design purposes. Many design factors must be considered for actual plant design.
    2. Dependent upon acid gas partial pressures and solution concentrations.
    3. Dependent upon acid gas partial pressures and corrosiveness of solution. Might be only 60% or less of value shown for corrosive systems.
    4. Varies with stripper overhead reflux ratio. Low residual acid gas contents require more stripper trays and/or higher reflux ratios yielding larger reboiler duties.
    5. Varies with stripper overhead reflux ratios, rich solution feed temperature to stripper and reboiler temperature.
    6. Maximum point heat flux can reach 20,00025,000 Btu/hr-ft2 at highest flame temperature at the inlet of a direct fired fire tube. The most satisfactory design of firetube heating elements employs a zone by zone calculation based on thermal efficiency desired and limiting the maximum tube wall temperature as required by the solution to prevent thermal degradation. The average heat flux, Q/A, is a result of these calculations.
    7. Reclaimers are not used in DEA and MDEA systems.
    8. Reboiler temperatures are dependent on solution conc. flare/vent line back pressure and/or residual CO2 content required. It is good practice to operate the reboiler at as low a temperature as possible. Process Description
    A typical amine system is shown in Figure 6.3. Sour gas enters the system through an inlet separator (scrubber) to remove any entrained water or hydrocarbon liquids. Gas enters the bottom of the amine absorber (contactor) and flows upward, countercurrent to the lean amine solution which flows down from the top. The lean amine that returns to the contactor is maintained at a temperature above the vapor that exits the contactor to prevent any condensation of heavier liquid hydrocarbons. Intimate contact between the gas and amine solution is achieved by use of either trays or packing in the contactor.
    The absorber tower consists of trays (diameters >20 in. (500 mm)), conventional packing (diameters <20 in.) or structured packing (diameters >20 in.).
    Sweetened gas leaves the top of the tower. An optional outlet (separator) scrubber may be included to recover entrained amine from the sweet gas. Since the natural gas leaving the top of the tower is saturated with water, the gas will require dehydration before entering a pipeline.
    Rich amine, solution containing CO2 and H2S, leaves the bottom of the absorber and flows to the flash tank (drum) where most of the dissolved hydrocarbon gases or entrained hydrocarbon condensates are removed. A small amount of the acid gases flash to the vapor phase. From the flash drum, the rich amine proceeds to the rich amine/lean amine heat exchanger where it recovers some of the sensible heat from the lean amine stream, which decreases the heat duty on the amine reboiler and the solvent cooler. Preheated rich amine then enters the amine stripping tower where heat from the reboiler breaks the bonds between the amine and acid gases. Acid gases are removed overhead and lean amine is removed from the bottom of the stripper.
    Hot lean amine flows to the rich amine/lean amine heat exchanger and then to additional coolers, typically aerial coolers, to lower its temperature about 100F (5.50C) above the inlet gas temperature. This reduces the amount of hydrocarbons condensed in the amine solution when the amine contacts the sour gas.
    A side stream of amine, of about 3%, is taken off after the rich/lean amine heat exchanger and is flowed through a charcoal filter to clean the solution of contaminants (not included in figure). Cooled lean amine is then pumped up to the absorber pressure and enters the top of the absorber. Amine solution flows down the absorber where it absorbs the acid gases. Rich amine is then removed at the bottom of the tower and the cycle is repeated.

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    Chapter 6 - Part 1 B General Remarks
    The contactor operates above ambient temperature because of the combined exothermic heat of absorption and reaction. The maximum temperature is in the lower portion of the tower because the majority of the absorption and reaction occurs near the bottom of the unit. The temperature bulge in the tower can be up to about 180F (80C). The treated gas leaves the top of the tower water saturated and at a temperature controlled by the temperature of the lean amine that enters, usually around 100F (38C).
    The sweet gas leaving the contactor is saturated with water so dehydration, is normally required prior to sale. If MEA is the sweetening agent, or the contactor is operating at unusually high temperature, a water wash may be used to attempt to recover some of the vaporized and/or entrained amine from the gas leaving the contactor. If a water wash is used it generally will consist of three or four trays at the top of the contactor, with makeup water to the unit being used as the wash liquid.
    Acid gas stripped from the amine passes out of the top of the stripper. It goes through a condenser and separator to cool the stream and recover water. The recovered water is usually returned to the stripper as reflux. The acid gas from the reflux separator is either vented, incinerated, sent to sulfur recovery facilities, or compressed for sale or reinjected into a suitable reservoir for enhanced oil recovery projects or for sequestration.

    Fig. 6-3. Gas Sweetening by Amine absorption, operating conditions are representative, not definitive. Operating Issues
    CorrosionSome of the major factors that affect corrosion are:
    Amine concentration (higher concentrations favor corrosion)
    Rich amine acid gas loading (higher gas loadings in the amine favor corrosion)
    Oxygen concentration
    Heat stable salts (higher concentrations promote corrosion and foaming)
    In addition to destroying vessels and piping, the corrosion products can cause foaming.

    Solution FoamingFoaming of the liquid amine solution is a major problem because it results in poor vapor−liquid contact, poor solution distribution, and solution holdup with resulting carryover and off spec gas. Among the causes of foaming are suspended solids, liquid hydrocarbons, surface active agents, such as those contained in inhibitors and compressor oils, and amine degradation products, including heat stable salts. One obvious cure is to remove the offending materials; the other is to add antifoaming agents.

    Heat Stable Salts (HSS) As mentioned above, these amine degradation products can cause both corrosion and foaming. They are normally dealt with through the use of amine reclaimers. Reclaimer
    A reclaimer is usually required for MEA and DGA amine-based systems. The reclaimer helps remove degradation products from the solution and also aids in the removal of heat stable salts, suspended solids, acids and iron compounds. The reclaimers in MEA and DGA systems differ. For MEA, a basic solution helps reverse the reactions. Soda ash and/or caustic soda is added to the MEA reclaimer to provide a pH of approximately 8-9; no addition is required for the DGA reclaimer system. Reclaimers generally operate on a side stream of 1-3% of the total amine circulation rate. Reclaimer sizing depends on the total inventory of the plant and the rate of degradation expected.
    Reclaimer operation is a semi-continuous batch operation. The reclaimer is filled with hot amine solution and, if necessary, soda ash is added. As the temperature in the reclaimer increases, the liquid will begin to distill. Overhead vapors can be condensed and pumped back into the amine system, but generally the reclaimer is operated at slightly above stripper column pressure and the vapors are returned to the stripper.
    The initial vapor composition is essentially water. Continued distillation will cause the solution to become more and more concentrated with amine. This raises the boiling point of the solution and amine will begin to distill overhead. Fresh feed is continually added until the boiling point of the material in the reclaimer reboiler reaches 280 to 300F. At this point, distillation is continued for a short time adding only water to help recover residual amine in the reclaimer reboiler. The reclaimer is then cleaned, recharged, and the cycle is repeated.
    Reclaimer "sludge" removed during cleaning must be handled with care. Disposal of the "sludge" must be in accordance with the governing regulations. If needed a reclaiming company may be contracted to remove degradation products or heat stable salts from the amine.
    DEA does not form a significant amount of nonregenerable degradation products, and it requires more difficult reclaiming through vacuum distillation or ion exchange. Basic Amine Chemistry
    Amines are bases, and the important reaction in gas processing is the ability of the amine to form salts with the weak acids formed by H2S and CO2 in an aqueous solution. When a gas stream that contains the H2S, CO2, or both, is contacted by a primary or secondary amine solution, the acid gases react to form a soluble acid−base complex, a salt, in the treating solution. The reaction between the amine and both H2S and CO2 is highly exothermic. Regardless of the structure of the amine, the overall reactions between H2S and amines are simple since H2S reacts directly and rapidly with all amines to form the bisulfide and the sulfide by Eqs. 6-1 & 6-2:

    For hydrogen sulfide removal
    RNH2 + H2S ↔ RNH3+ + HS Fast Eq. 6-1
    RNH2 + HS ↔ RNH3+ + S2 Fast Eq. 6-2

    For carbon dioxide removal
    2RNH2 + CO2 ↔ RNH3+ + RNHCOO Fast Eq. 6-3
    RNH2 + CO2 + H2O ↔ RNH3+ + HCO3 Slow Eq. 6-4
    RNH2 + HCO3 ↔ RNH3+ + CO32 Slow Eq. 6-5

    Concerning the chemical reactions with CO2, primary amines (RNH2) such as MEA and DGA agent, and secondary amines (RRNH) such as DEA and DIPA, differ from tertiary amines (RRR"N) such as TEA and MDEA.
    For the reactions discussed above, high pressures and low temperatures drive the reactions to the right, whereas high temperatures and low pressures favor the reverse reaction, which thus provides a mechanism for regeneration of the amine solution.

    Primary and Secondary Amines
    With the primary and secondary amines, the predominant overall reaction (Eq. 6-3) rapidly leads to the formation of a stable carbamate which is slow to further hydrolyze to bicarbonate.
    The other overall reactions leading to bicarbonate (Eq. 6-4) and to carbonate (Eq. 6-5) are slow because they have to proceed through the hydration of CO2.
    Therefore, according to Eq. 6-3 there is a theoretical limit to the chemical loading capacity of the primary and secondary amine solutions to 0.5 mole CO2 per mole of amine, even at relatively high partial pressures of CO2 in the gas to be treated.
    For primary and secondary amines, little difference exists between the H2S and CO2 reaction rates because of the availability of the rapid carbamate formation for CO2 absorption. Therefore, the primary and secondary amines achieve essentially complete removal of H2S and CO2.

    Tertiary Amines
    Unlike primary and secondary amines, the nitrogen (N) in tertiary amines (RRR"N) has no free hydrogen ( H ) to rapidly form carbamate as per overall Eq. 6-3. Consequently, the removal of CO2 by tertiary amines can only follow the slow route to bicarbonate by Eq. 6-4 and carbonate by Eq. 6-5.
    The slowness of the reaction leading to bicarbonate is the underlying reason why tertiary amines can be considered selective for H2S removal, by playing with absorption contact time, and this attribute can be used to full advantage when complete CO2 removal is not necessary.
    However, the slow route to bicarbonates theoretically allows at equilibrium a chemical loading ratio of one mole of CO2 per mole of amine. Furthermore, at high partial pressure, the solubility of CO2 in tertiary amines is far greater than in the primary and secondary amines thus further enhancing the CO2 loading by physical solubility at high partial pressure.
    Therefore, in case of gases to be treated for bulk CO2 removal, large amounts of CO2 can be liberated from the rich solvent by simple flash alleviating the thermal regeneration duty with consequent energy savings. In other words, because the tertiary amines have no labile hydrogen, they cannot form the carbamate. Tertiary amines must react with CO2 via the slow hydrolysis mechanism in Equation 6.4. With only the slow acid−base reaction available for CO2 absorption, MDEA (methyldiethanolamine) and several of the formulated MDEA products yield significant selectivity toward H2S relative to CO2, and, consequently, all of the H2S is removed while some of the CO2 slips through with the gas.

    Activated Tertiary Amines
    The use of activators mitigates the slowness of the reaction to bicarbonate for tertiary amines. Activators are generally primary or secondary amines; they are tailored to increase both the hydrolysis of the carbamate and the rate of hydration of dissolved CO2 thus making the activated-tertiary amines specially suitable for efficient and economical bulk CO2 removal when selectivity is not required (see section on MDEA). Amines Used
    MEA is a primary amine, which has had widespread use as a gas sweetening agent. The process is well proven and can meet pipeline specifications. MEA is a stable compound and, in the absence of other chemicals, suffers no degradation or decomposition at temperatures up to its normal boiling point.
    Gas sweetening with monoethanolamine (MEA) is used where there are low contactor pressures and/or stringent acid gas specifications. MEA removes both H2S and CO2 from gas streams. H2S concentrations well below 4.0 ppmv can be achieved. CO2 concentrations as low as 100 ppmv can be obtained at low to moderate pressures.
    COS and CS2 are removed by MEA, but the reactions are irreversible unless a reclaimer is used. Even with a reclaimer, complete reversal of the reactions may not be achieved. The result is solution loss and build-up of degradation products in the system. Total acid gas pick up is traditionally limited to 0.3-0.35 moles of acid gas/mole of MEA and solution concentration is usually limited to 10-20 wt%. Inhibitors can be used to allow much higher solution strengths and acid gas loadings. Because MEA has the highest vapor pressure of the amines used for gas treating, solution losses through vaporization from the contactor and stripper can be high (MEA losses of 1-3 lbs/MMscf (16-48 kg/ MM m3) of inlet gas). This problem can be minimized by using a water wash.
    MEA reactions are reversible by changing the system temperature. Reactions with CO2 and H2S are reversed in the stripping column by heating the rich MEA to about 2450F at 10 psig (118 0C at 69 kPa). Acid gases evolve into the vapor and are removed from the still overhead. Thus, the MEA is regenerated.
    MEA systems foam rather easily resulting in excessive amine carryover from the absorber. Foaming can be caused by a number of foreign materials such as condensed hydrocarbons, degradation products, solids such as carbon or iron sulfide, excess corrosion inhibitor, and valve grease.
    A microfiber filter separator should be installed at the gas inlet to the MEA contactor. It is an effective method of foam control and removes many of the contaminants before they enter the system. Hydrocarbon liquids are usually removed in the flash tank. Degradation products are removed in a reclaimer as described above.
    A gas blanket system is installed on MEA storage tanks and surge vessels. This prevents oxidation of MEA. Sweet natural gas or nitrogen is normally used.
    Disadvantages include the reaction of MEA with carbonyl sulfide (COS) and carbon disulfide (CS2) to form heat-stable salts, which cannot be regenerated at normal stripping column temperatures. At temperatures above 245 0F (118 0C) a side reaction with CO2 exists that produces oxazolidone-2, a heat-stable salt, which consumes MEA from the process. Normal regeneration temperature in the still will not regenerate heat-stable salts or oxazolidone-2. A reclaimer is often included to remove these contaminants.

    This process employs an aqueous solution of diethanolamine (DEA) (a secondary amine). DEA will not treat to pipeline quality gas specifications at as low a pressure as will MEA.
    This process is used for high pressure, high acid gas content streams having a relatively high ratio of H2S/CO2.
    Although mole/mole loadings as high as 0.8-0.9 have been reported, most conventional DEA plants still operate at significantly lower loadings.
    The process flow scheme for conventional DEA plants resembles the MEA process. The advantages and disadvantages of DEA as compared to MEA are:

    The mole/mole loadings typically used with DEA (0.35-0.82 mole/mole) are much higher than those normally used (0.3-0.4) for MEA.
    Molecular weight of DEA is 105 compared to 61 for MEA. Combination of molecular weights and reaction stoichiometry means that about 1.7 lbs (0.77 kg) of DEA must be circulated to react with the same amount of acid gas as 1.0 lbs (0.45 kg) of MEA. The solution strength of DEA ranges up to 35% by weight compared to 20% for MEA. Loadings for DEA systems range from 0.35 to 0.65 mol of acid gas per mole of DEA without excessive corrosion. The result of this is that the circulation rate of a DEA solution is slightly less than in a comparable MEA system.
    Because DEA does not form a significant amount of nonregenerable degradation products, a reclaimer is not required.
    DEA is a secondary amine and is chemically weaker than MEA, and less heat is required to strip the amine solution.
    DEA forms a regenerable compound with COS and CS2 and can be used for the partial removal of COS and CS2 without significant solution losses.
    As a secondary amine, DEA is less alkaline than MEA. DEA systems do suffer the same corrosion problems, but not as severely as those using MEA. Solution strengths are typically from 25% to 35% DEA by weight in water.
    Vapor pressure of DEA is about 1/30 of the vapor pressure of MEA. Thus, DEA amine losses are much lower than in an MEA system.

    DGA is a primary amine capable of removing not only H2S and CO2, but also COS and mercaptans from gas and liquid streams. DGA has been used to treat natural gas to 4.0 ppmv at pressures as low as 125 psig. Compared with MEA, low vapor pressure allows Diglycolamine [ 2-(2-aminoethoxy) ethanol] (DGA) is reclaimed on-site to remove heat stable salts and reaction products with COS and CS2.
    DGA has a greater affinity for the absorption of aromatics, olefins, and heavy hydrocarbons than the MEA and DEA systems.
    Therefore, adequate carbon filtration should be included in the design of a DGA treating unit.
    The process flow for the DGA treating process is similar to that of the MEA treating process. The major differences are:
    Higher acid gas pick-up per gallon of amine can be obtained by using 50-60% solution strength rather than 15-20% for MEA (more moles of amine per volume of solution).
    The required treating circulation rate is lower. This is a direct function of higher amine concentration.
    Reduced reboiler steam consumption.
    DGA has an advantage for plants operating in cold climates where freezing of the solution could occur. The freezing point for 50% DGA solution is 30F.
    Low vapor pressure decreases amine losses.
    Unlike MEA, degradation products from reactions with COS and CS2 can be regenerated.

    Methyldiethanolamine (MDEA) is a tertiary amine which can be used to selectively remove
    H2S to pipeline specifications at moderate to high pressure. If increased concentration of CO2 in the residue gas does cause a problem with contract specifications or downstream processing, further treatment will be required.
    If the gas is contacted at pressures ranging from 800 to 1000 psig, H2S levels can be reduced to concentrations required by pipelines. While at the same time, 40-60% of the CO2 present flows through the contactor, untreated.
    The H2S/CO2 ratio in the acid gas can be 10-15 times as great as the H2S/CO2 ratio in the sour gas. Some of the benefits of selective removal of H2S include:
    Reduced solution flow rates resulting from a reduction in the amount of acid gas removed.
    Smaller amine regeneration unit. Significant capital savings are realized due to reduced pump and regeneration requirements.
    Higher H2S concentrations in the acid gas resulting in reduced problems in sulfur recovery.
    MDEA has a lower heat requirement due to its low heat of regeneration. In some applications, energy requirements for gas treating can be reduced as much as 75% by changing from DEA to MDEA.
    It is not reclaimable by conventional methods.
    CO2 hydrolyzes much slower than H2S. This makes possible significant selectivity of tertiary amines for H2S. This fact is used for selective removal of H2S from gases containing both H2S and CO2.
    A feature of MDEA is that it can be partially regenerated in a simple flash. As a consequence the removal of bulk H2S and CO2 may be achieved with a modest heat input for regeneration.
    However, as MDEA solutions react only slowly with CO2, activators must be added to the MDEA solution to enhance CO2 absorption and the solvent is then called activated MDEA.
    Solution strengths typically range from 40% to 50% MDEA by weight. Acid gas loading varies from 0.2 to 0.4 or more moles of acid gas per mole of MDEA depending on the supplier.

    Diisopropanolamine (DIPA) is a secondary amine which exhibits, though not as great as tertiary amines, selectivity for H2S. It is similar to DEA systems but offers the following advantages:
    Carbonyl sulfide (COS) can be removed and the DIPA solution can be regenerated easily
    The system is generally noncorrosive and has a lower energy consumption.
    At low pressures, DIPA will preferentially remove H2S. As pressure increases, the selectivity of the process decreases and DIPA removes increasing amounts of CO2. Thus, this system can be used either to selectively remove H2S or to remove both CO2 and H2S.

    Formulated Solvents and Mixed Amines
    The selectivity of MDEA can be reduced by addition of MEA, DEA, or proprietary additives. Thus, it can be tailored to meet the desired amount of CO2 slippage and still have lower energy requirements than do primary and secondary amines.
    Formulated Solvents is the name given to a new family of amine based solvents. Their popularity is primarily due to equipment size reduction and energy savings over most of the other amines. All the advantages of MDEA are valid for the Formulated Solvents, usually to a greater degree. Some formulations are capable of slipping larger portions of inlet CO2 (than MDEA) to the outlet gas and at the same time removing H2S to less than 4 ppmv. For example, under conditions of low absorber pressure and high CO2 /H2S ratios, such as Claus tail gas clean-up units, certain solvent formulations can slip upwards to 90 percent of the incoming CO2 to the incinerator.
    While at the other extreme, certain formulations remove CO2 to a level suitable for cryogenic plant feed. Formulations are also available for CO2 removal in ammonia plants. Finally, there are solvent formulations which produce H2S to 4 ppmv pipeline specifications, while reducing high inlet CO2 concentrations to 2% for delivery to a pipeline. This case is sometimes referred to as bulk CO2 removal.
    This need for a wide performance spectrum has led Formulated Solvent suppliers to develop a large stable of different MDEA-based solvent formulations. Most Formulated Solvents are enhancements to MDEA discussed above. Thus, they are referred to as MDEA-based solvents or formulations.
    Benefits claimed by suppliers are:
    For New Plants
    reduced corrosion
    reduced circulation rate
    lower energy requirements
    smaller equipment due to reduced circulation rates
    For Existing Plants
    increase in capacity, i.e., gas through-put or higher inlet acid gas composition
    reduced corrosion
    lower energy requirements and reduced circulation rates

    Inhibited Amine Systems
    Inhibited amine processes use standard amines that have been combined with special inhibiting agents that minimize corrosion. They allow higher solution concentrations and higher acid gas loadings, thus reducing required circulation rates and energy requirements. They also utilize hot potassium carbonate to remove CO2 and H2S. As a general rule, this process should be considered when the partial pressure of the acid gas is 20 psia (138 kPa) or greater. The process is not recommended for low-pressure absorption, or high-pressure absorption of low-concentration acid gas.

    Sterically Hindered Amines
    In acid gas removal, steric hindrance involves alteration of the reactivity of a primary or secondary amine by a change in the alkanol structure of the amine.
    A large hydrocarbon group attached to the nitrogen shields the nitrogen atom and hinders (inhibits) the carbamate reaction. The H2S reaction is not significantly affected by amine structure, because the proton is small and can reach the nitrogen. However, CO2 removal can be significantly affected if the amine structure hinders the fast carbamate formation reaction and allows only the much slower bicarbonate formation.

    Table 6.7. Some Representative Operating Parameters for Amine Systems Heats of Reaction
    Amines are basic solutions. These solutions react with the hydrogen sulfide and carbon dioxide to form a salt. The process of absorbing the acid gases generates heat.
    The magnitude of the exothermic heats of reaction, which includes the heat of solution, of the amines with the acid gases is important because the heat liberated in the reaction must be added back in the regeneration step. Thus, a low heat of reaction translates into smaller energy regeneration requirements. Table 6.8 summarizes the important data for the common amines. The values in Table 6.8 are approximate because heats of reaction vary with acid gas loading and solution concentration.

    Table 6.8. Average Heats of Reaction of the Acid Gases in Amine Solutions

    6.3.2 Alkaline Salt Process (Hot Carbonate)
    The hot potassium carbonate process for removing CO2 and H2S was developed by the United States Bureau of Mines and is described by Benson and coworkers in two papers. Although the process was developed for the removal of CO2, it can also remove H2S if H2S is present with CO2 (not suitable for gas streams containing only H2S.). Special designs are required for removing H2S to pipeline specifications or to reduce CO2 to low levels.
    The process is very similar in concept to the amine process, in that after physical absorption into the liquid, the CO2 and H2S react chemically with the solution. The chemistry is relatively complex, but the overall reactions are represented by

    K2CO3 + CO2 + H2O ↔ 2KHCO3 Eq. 6-6
    K2CO3 + H2S ↔ KHS + KHCO3 Eq. 6-7

    In a typical application, the contactor will operate at approximately 300 psig (20 barg), with the lean carbonate solution entering near 225F (110C) and leaving at 240F (115C). The rich carbonate pressure is reduced to approximately 5 psig (0.3 barg) as it enters the stripper.
    Approximately one third to two thirds of the absorbed CO2 is released by the pressure reduction, reducing the amount of steam required for stripping. The lean carbonate solution leaves the stripper at the same temperature as it enters the contactor, and eliminates the need for heat exchange between the rich and lean streams. The heat of solution for absorption of CO2 in potassium carbonate is small, approximately 32 Btu/cu ft of CO2, and consequently the temperature rise in the contactor is small and less energy is required for regeneration.
    Potassium carbonate processes are somewhat effective in removing carbonyl sulfide and carbon disulfide. Potassium carbonate works best on gas streams with a CO2 partial pressure of 30-90 psi.
    Pipeline-quality gas often requires secondary treating using an amine or similar system to reduce H2S level to 4 ppm.
    Because this system is operated at high temperatures to increase the solubility of carbonates, the designer must be careful to avoid dead spots in the system where the solution could cool and precipitate solids. If solids do precipitate, the system may suffer from plugging, erosion, or foaming.
    Hot potassium carbonate solutions are corrosive. All carbon steel must be stress relieved to limit corrosion. Varieties of corrosion inhibitors, such as fatty amines or potassium dichromate, are available to decrease corrosion rates. There are three basic process flow variations for the potassium carbonate process. The flow scheme required depends on the outlet specification of the natural gas. These are:

    Single Stage Process
    The single stage process is shown in Fig. 6-4. Potassium carbonate is pumped to the top of a packed or trayed contactor where it contacts the gas stream. The rich solution flows to the stripper where the acid gases are stripped with steam. The lean solution is then pumped back to the contactor to complete the cycle.

    Split Flow Process
    In this process scheme (Fig. 6-5) the lean solution stream is split. Hot solution is fed to the middle of the contactor for bulk removal. The remainder is cooled to improve equilibrium and is fed to the top of the contactor for trim acid gas removal.

    Two Stage Process
    In this process scheme (Fig. 6-6) the contactor is like that of the split flow process. In addition, the stripper is in two sections. A major portion of the solution is removed at the midpoint of the stripper and pumped to the lower section of the contactor.
    Numerous improvements have been made to the potassium carbonate process resulting in significant reduction in capital and operating costs. At the same time, lower acid gas concentration in the treated gas can now be achieved. The most popular of the carbonate processes are:

    Benfield Process
    The Benfield Process is licensed by UOP. Several activators are used to enhance the performance of the potassium carbonate solution.

    Hi-Pure Process
    The Hi-Pure process is a combination conventional Benfield potassium carbonate process and alkanolamine process. The gas stream is first contacted with potassium carbonate followed by contacting with an amine. The process can achieve outlet CO2 concentrations as low as 30 ppmv and H2S concentrations of 1 ppmv.

    Catacarb Process
    The Catacarb Process is licensed by Eickmeyer and Associates. Activators, corrosion inhibitors, potassium salts, and water are contained in the solution. This process is mostly used in the ammonia industry.

    Fig. 6-4. Alkaline Salt: Single-Stage Process

    Fig. 6-5. Alkaline Salt: Split-Flow Process

    Fig. 6-6. Alkaline Salt: Two-Stage Process

    6.3.3 Specialty Batch Chemical Solvents
    Several batch chemical processes have been developed and have specific areas of application.
    Processes include

    General Process Description
    Gas is flowed into a vessel and contacted with the solvent. Acid components are converted to soluble salts, which are nonregenerable, limiting the life of the solution.
    Once saturation levels are reached, the solution must be replaced.
    For some of these processes, the spent solutions are not hazardous, but for others, the spent solutions have been labeled hazardous and, if used, must be disposed of as Class IV materials.
    Units in these processes have a wide operating range, with acid gas concentrations ranging from as low as 10 ppm to as high as 20%. Operating pressures range from near atmospheric to >1000 psig (7000 kPa). Some units have been designed to handle from several thousand cubic feet per day to more than 15 MMscfd (several hundred cubic meters per day to more than 420,000 m3 per day).

    a- Sulfa-Check
    Sulfa-Check is a product of ExxonMobil Chemicals which selectively removes H2S and mercaptans from natural gas in the presence of CO2. The patented process converts the sour gas directly to sulfur. This is accomplished by sparging bubbling the gas in a buffered, water based oxidizing solution containing sodium nitrite (NaNO2). The sodium nitrite is reduced to ammonia (NH3) which remains in solution. The spent product is classified as non-hazardous.
    Dozens of field applications include gas rates ranging from 15 Mscf/d to 3.0 MMscf/d and inlet H2S concentrations ranging from 10 to 3000 ppmv.
    Reaction rate is independent of the concentration of the oxidizing agent.
    There is no limit to the concentration of H2S treated. Process is most economical for acid gas streams containing from1 ppm to 1% H2S. pH must be held above 7.5 to control selectivity and optimize H2S removal. One gallon (4 l) of oxidizing solution can remove up to 2 lbs (1 kg) of H2S when the system is operated at ambient temperatures<100 0F (38 0C). If gas temperatures exceed 100 0F (38 0C), the solubility of sulfur in the oxidizing agent decreases.
    Operating pressure of at least 20 psig is required for proper unit operation to maintain bubble flow through the column. Bubble flow is necessary to produce intimate mixing of the gas and liquid.
    Oxidizing solution will eventually become saturated and require replacement. Disposal of this slurry poses no environmental problem, as the reaction produces an aqueous slurry of sulfur and sodium salt.
    A number of variables, including some associated risks must be considered prior to determining if the Sulfa-Check process is applicable. For example, low levels of ammonia may appear in the treated gas. Also, the reduction of NO2 may result in the formation of nitric oxide (NO). If air is present in the raw gas, it will react with nitric oxide to form nitrogen dioxide (NO2). NO2 is a strong oxidizing agent that will react with elastomers and odorants and cause corrosion in a moist environment.

    b- Caustic Wash
    Caustic (NaOH) scrubbing systems can be used to treat natural gas streams to remove CO2, CS2, H2S, and mercaptans. The process employs countercurrent contacting of the gas stream with a caustic solution in a packed or trayed column. The column may contain one stage or several stages depending on the required degree of removal. The multi-stage systems generally have different caustic concentrations ranging from 4-6 weight percent in the first stage to 8-10 weight percent in the latter stages. Multiple stages increase the caustic efficiency while maintaining a sufficient driving force to achieve absorption.
    The spent solution is either regenerated or discarded depending on what acid gas components are present in the gas stream. If only mercaptans are present, the caustic solution is regenerated with steam in a stripping still. If CO2 is present, a nonregenerable product (Na2CO3) is formed and the solution must be discarded. As a result, the presence of CO2 in caustic systems leads to high caustic consumption. This is a serious disadvantage of the caustic scrubbing process. The spent caustic solutions are considered hazardous wastes. Natural gas is usually water washed after a caustic wash to remove any caustic entrained in the gas prior to dehydration.

    The chemical reactions involved are as follows:
    H2S + 2 NaOH → Na2S + 2H2O Eq 6-8
    RSH + NaOH → RSNa + H2O Eq 6-9
    CO2 + 2 NaOH → Na2CO3 + H2O Eq 6-10
    CS2 + 2 NaOH → 2 NaHS + CO2 Eq 6-11

    Fig. 6-7. Regenerative Caustic

    c- Sulfa-Scrub
    Sulfa-Scrub from Baker-Petrolite is a product that uses triazine compound to selectively react with H2S. The raw gas can be sparged into a vessel containing the liquid scavenging agent. Alternatively the liquid may be injected to the gas stream as an H2S scavenger. The spent material is considered as non-hazardous and is claimed to be an excellent corrosion inhibitor.
    6.4 Physical Solvent Processes
    In the amine and alkali salt processes, the acid gases are removed in two steps:
    physical absorption followed by chemical reaction. In processes such as Selexol or Rectisol , no chemical reaction occurs and acid gas removal depends entirely on physical absorption. Some of the inherent advantages and disadvantages of physical absorption processes are summarized below:
    These processes flow scheme is as shown in Fig. 6-8. In general, a physical solvent process should be considered when:
    1- The partial pressure of the acid gas in the feed is greater than 50 psi.
    2- The heavy hydrocarbon concentration in the feed gas is low.
    3- Bulk removal of the acid gas is desired.
    4- Selective removal of H2S is desired.

    These processes are economically attractive because little energy is required for regeneration. The solvents are regenerated by:
    1- Multi-stage flashing to low pressures.
    2- Regeneration at low temperatures with an inert stripping gas.
    3- Heating and stripping of solution with steam/solvent vapors.

    In general, physical solvents are capable of removing COS, CS2, and mercaptans.
    In certain instances, physical absorption processes are capable of simultaneously dehydrating and treating the gas although additional equipment and higher energy requirements may be needed to dry the solvent. The processes operate at ambient or subambient temperature to enhance the solubility of the acid gases. The solvents are relatively noncorrosive so carbon steel can be used. Chemical losses are low due to low solvent vapor pressure or refrigerated conditions. Physical solvents will absorb heavy hydrocarbons from the gas stream resulting in high hydrocarbon content in the acid gas stream as well as possibly significant hydrocarbon losses. Some of the physical absorption processes are summarized below.

    Fig. 6-8. Typical Gas Sweetening by Physical Absorption

    6.4.1 Selexol
    This process developed by Allied Chemical Corp. uses a polyethylene glycol derivative as a solvent. The solvent is selective for RSH, CS2, H2S, and other sulfur compounds. The process can be used to selectively or simultaneously remove sulfur compounds, carbon dioxide, water, as well as paraffinic, olefinic, aromatic and chlorinated hydrocarbons from a gas or air stream. Because water and heavy hydrocarbons are highly soluble in Selexol, the treated gas from a Selexol unit normally meets both water and hydrocarbon dew point specifications.
    Levels of CO2 can be reduced by about 85%. The process is economical when high acid gas partial pressures exist and there is an absence of heavy ends in the gas.
    This process will not typically remove enough CO2 to meet pipeline gas requirements. DIPA can be added to the solution to remove CO2 down to pipeline specifications. The process also removes water to <7 lbs/MMscf (0.11 g/std m3). The addition of DIPA increases the relatively low stripper heat duty.
    The vendor states that the solvent is very stable, no degradation products are formed or disposed of, and no solvent reclaiming is required. Depending on the applications, the operating pressure could be as low as ambient though higher pressure is preferred. Operating temperature varies from 0F to ambient. Selexol is a Dow Chemical solvent and a UOP technology.
    Table 6-9 presents RK, the ratio of the K-value* for methane, KCH4 (arbitrarily assigned a value of 1), to the K-values of the other component, RK = KCH4/Kcomponent . (The K-value is the ratio of the mole fraction of the component in the vapor phase (y) to its mole fraction in the liquid phase (x), K = y/x. High K-values indicate the material is predominately in the vapor phase, whereas low K-values indicate a higher concentration in the liquid phase (x)).

    Table 6-9 Typical Relative Ratio of K-values

    An R value greater than1 indicates that the solubility of the component in Selexol is greater than that of methane. The values should be regarded as only representative because pressure and temperature are not specified and, as previously noted, the composition of Selexol is variable.
    Because RK for CO2 and H2S are 15 and 134, respectively, these gases are preferentially absorbed (relative to CH4), and, consequently, physical absorption is an effective technique for acid gas removal. The process can reduce H2S to 4 ppmv, reduce CO2 to levels below 50 ppmv, and essentially remove all mercaptans, CS2, and COS. Two additional features of Table 6-9 are worth mentioning. Because the RK values for hydrocarbons heavier than CH4 are fairly high (6.4 for C2H6, 15.3 for C3H8, and 35 for n-C4H10), Selexol will remove substantial quantities of these hydrocarbons, a feature that can be either positive or negative, depending on the composition of the gas being processed and the desired products. Finally, the RK value of H2O is extremely high and consequently, Selexol provides some dehydration.
    Volumes of some gases in scf absorbed/gal Selexol are plotted in Figure 6-9. The figure assumes a Henrys law relationship, (For an ideal system, Henrys law assumes a linear relation between the solubility of gas component (i) and its partial pressure, yiP = kixi where ki is the Henrys constant.) which provides approximate solubility at higher pressures. The lines also ignore probable Interaction between solutes.
    Applications of Selexol are varied and, consequently, no common process flow diagrams are available. One plant is shown in Figure 6-10. The plant pretreats the gas to reduce CO2, ethane, and heavier hydrocarbon levels before final purification in molecular sieve units and subsequent liquefaction. The plant is designed to process 26 MMscfd (0.74 MMSm3/d) entering the Selexol unit at 603 psia (41.6 bar) and 32F (0C). The lean solvent, cooled to 25F (−3.9C) with propane refrigerant, enters the absorber where it absorbs CO2 and some of the ethane and heavier hydrocarbons. The rich solvent from the absorber is regenerated by reduction of the pressure in three flash drums, from 603 to 106 psia (41.6 to 7.3 bar) in the high-pressure drum, from 106 to 16 psia (7.3 to 1.1 bar) in the medium-pressure drum, and from 16 to 3 psia (1.1 to 0.21 bar) in the vacuum drum. Lean Selexol from the vacuum drum is recompressed and sent to the propane chiller. The treated gas that leaves the absorber passes through a knockout drum and filter separator to remove entrained Selexol and condensed hydrocarbons. Table 6-11 shows that the treated gas meets the specifications of a maximum of 0.50% CO2 and a maximum of 6.5% ethane and heavier hydrocarbons. In addition, the water content of the gas is reduced from 75 ppmv to 12 ppmv, H2S is reduced from 2 ppmv to essentially nothing, and methyl mercaptan is reduced from 5 ppmv to 1 ppmv. Unlike the amine systems, no irreversible products are generated in the process, which thus eliminates the need for reclaiming.

    Table 6-10. Representative Property Data for Selexol

    Fig. 6-9. Solubility of various gases in Selexol solvent at 70F (21C) as a function of partial pressure.

    Table 6-11. Composition of Inlet and Outlet Gas in a Selexol Unit

    Fig. 6-10. Process schematic for a Selexol gas treating facility.

    6.4.2 Fluor Solvent
    The Fluor Solvent process uses propylene carbonate as a physical solvent to remove CO2 and H2S. Propylene carbonate also removes C3+ hydrocarbons, COS, SO2, CS2, and H2O from the natural gas stream. Thus, in one step the natural gas can be sweetened and dehydrated to pipeline quality.
    This process is used for bulk removal of CO2 and is not used to treat to <3% CO2.
    This system requires special design features such as larger absorbers and higher circulation rates to obtain pipeline quality and usually is not economically applicable for outlet requirements.
    Propylene carbonate has the following characteristics, which make it suitable as a solvent for acid gas treating:
    High degree of solubility for CO2 and other gases
    Low heat of solution for CO2
    Low vapor pressure at operating temperature
    Low solubility for light hydrocarbons (C1, C2)
    Chemically nonreactive toward all natural gas components
    Low viscosity
    Noncorrosive toward common metals
    The above characteristics combine to yield a system that has low heat and pumping requirements, is relatively noncorrosive, and suffers only minimal solvent losses (<1 lbs/MMscf).
    Solvent temperatures below ambient are often used to increase solvent gas capacity and, therefore, decrease circulation rates. Expansion of the rich solvent and flash gases through power turbines can provide the required refrigeration. Alternately, auxiliary refrigeration may be included to further decrease circulation rates.

    6.4.3 Rectisol Process
    This process uses pure refrigerated methanol as a solvent and has been developed and licensed by Lurgi Oel Gas Chemie and Linde. The process is often applied for syngas purification and operates at temperatures as low as minus 30 to minus 100F. Applications include selective or bulk removal of sour components. However the extremely low total sulfur impurities achievable with this process (down to less than 0.1 ppmv), are normally not required in natural gas service. Here the process is best suited where there are very high amounts of CO2 to be removed.
    The process has been applied for the purification of natural gas for LNG production. However as pure cold methanol exhibits a certain co-absorption of higher hydrocarbons, implementation of the process in natural gas service has been limited to applications where only low concentrations of ethane and heavier components are present.

    6.4.4 Ifpexol Process
    This process developed by IFP licensed by Prosernat uses refrigerated methanol solutions containing water in order to reduce hydrocarbon co-absorption. It is generally associated downstream of the Ifpex-1 dehydration process (Dehydration chapter) which simultaneously extracts heavier hydrocarbons. The Ifpex-2 process attains equivalent acid gas and mercaptans removal performance of the Rectisol process but at a slightly higher solvent rate due to the lower solvent purity.

    6.4.5 Purisol
    This process has been developed and licensed by Lurgi Oel Gas Chemie. The solvent used is N-methyl-2-pyrrolidone (NMP or N-Pyrol), a high boiling point liquid. Purisol exhibits a selectivity for H2S, like Selexol, and features equivalent process possibilities.
    As Purisol is also well suited for the selective removal of mercaptans; it can be used for the recovery of mercaptans from regeneration off gases from adsorptive mercaptan removal units.

    6.4.5 Catasol
    This process is licensed by Eickmeyer and Associates. Catasol solvents are reported to have selectivity for H2S/CO2, CO2/propane, and COS/CO2.

    6.4.6 Morphysorb
    A physical solvent process developed by Krupp Uhde called Morphysorb, uses N-formyl morpholine (NFM) as a physical solvent. The process claims lower investment and operating costs when processing gases heavily loaded with acid components.

    6.5 Hybrid Solvent Processes
    Table 6-4 showed the strengths and weaknesses of amine and physical solvent systems. To take advantage of the strengths of each type, a number of hybrid processes commercially used, and under development, combine physical solvents with amines. Depending upon the solvent−amine combination, nearly complete removal of H2S, CO2, and COS is possible. Other hybrid systems provide high H2S and COS removal while slipping CO2. Sulfinol currently is one of the more commonly used processes. The process uses a combination of a physical solvent (sulfolane) with DIPA or MDEA. The selected amine depends upon the acid gases in the feed and whether CO2 removal is required. Like the physical solvent processes, the hybrid systems may absorb more hydrocarbons, including BTEX, but that property can be adjusted by varying water content.

    6.5.1 Sulfinol Process
    The Sulfinol Process, licensed by Shell Global Solutions, is used to remove H2S, CO2, COS, CS2, mercaptans and polysulfides from natural and synthetic gases. Sulfinol is a mixture of Sulfolane (a physical solvent), water and either DIPA or MDEA (both chemical solvents).
    Solution concentrations range between 25% and 40% sulfolane, 40-55% DIPA, and 20-30% water and depend on the conditions and composition of the gas being treated. It is this dual capacity as both a physical and a chemical solvent that gives Sulfinol its advantages.
    Sulfinol with DIPA (Sulfinol-D) is used when complete removal of H2S, CO2, and COS is desired. Sulfinol with MDEA (Sulfinol-M) is used for the selective removal of H2S in the presence of CO2, with partial removal of COS. Both Sulfinols can reduce the total sulfur content of treated gas down to low ppm levels. Some disadvantages are: a higher heavy hydrocarbon co-absorption, and a reclaimer is sometimes required for Sulfinol D when removing CO2.
    The presence of the physical solvent, Sulfolane@, allows higher acid gas loadings compared to systems based on amine only. Typical loadings are 1.5 mol of acid gas per mole of Sulfinol solution. Higher acid gas loadings, together with a lower energy of regeneration, can result in lower capital and energy costs per unit of acid gas removed as compared to the ethanolamine processes.
    Features of the Sulfinol process include essentially complete removal of mercaptans, high removal rate of COS, lower foaming tendency, lower corrosion rates, and the ability to slip up to 50% CO2.
    The design is similar to that of the ethanolamines. Degradation of DIPA to oxazolidones (DIPA-OX) usually necessitates the installation of a reclaimer for their removal.
    As with the ethanolamine processes, aromatics and heavy hydrocarbons in the feed gas should be removed prior to contact with the Sulfinol solution to minimize foaming.
    The merits of the Sulfinol process, as compared to the ethanolamine processes, are many, but there are other factors that must be considered before selecting the appropriate gas treating process. For example, a licensing fee, while not necessary for the ethanolamine processes, is required for the Sulfinol process. The solvent costs are generally higher for the Sulfinol process than they are for DEA. Operators are more familiar with DEA and the typical problems associated with this process. In cases of low acid gas partial pressure, the advantage of a lower circulation rate for the Sulfinol process diminishes compared to the DEA process.

    6.5.2 Selefining Process
    This process developed by Snamprogetti uses tertiary amines such as dimethyl ethanol amine (DMEA) dissolved in an organic solvent with limited amounts of water. The process is very selective to H2S as CO2 hydration is almost completely avoided. It also removes other sulfur species such as mercaptans, COS and CS2. It has a tendency to co-absorb hydrocarbons which can to some extent be controlled by increasing the water content of the solvent.

  6. Re: Basics of Gas Field Processing Book "Full text"

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

    6.6 General Considerations for Solvent Process
    6.6.1 Solution Filtration
    Filtration of the treating solution to remove entrained solids is essential to the successful operation of a gas treating plant.
    Filtration rates should be as high as practical and may range from 5 per cent of circulation to full stream. Removing particles down to 5 microns in size is recommended. In order to do this efficiently, two stages of filtration may be required. The first stage, typically a cartridge-type or precoat filter, is designed to remove particles down to the 10 micron or less range.
    The second stage of filtration, typically an activated carbon filter, removes hydrocarbons and other contaminants. This is accomplished by adsorption. The carbon filter can also remove smaller particles from the amine stream. The carbon granule size can be selected to remove particles down to the 5 micron range. The activated carbon filter should always be located downstream of the first stage filter because the deposition of solids would plug the carbon filter.
    The carryover of carbon fines can be controlled by either locating a second cartridge-type filter immediately downstream of the carbon filter or using a graded carbon bed. In a graded bed, larger granules are placed at the outlet of the filter to trap fines. Large carbon granules produce fewer fines but are less efficient for adsorption.
    Basic degradation products are identified by gas chromatography and mass spectrometry. Acidic degradation products are identified by ion chromatography exclusion. These tests are recommended when the amine solution appears to lose its ability to pick up acid gas. Degradation products affect the results of the conventional estimation of amine concentration by titration. This may cause artificially high or low apparent amine concentrations Also, the carbon bed will adsorb very little strong acid degradation products. In this case, purging or reclamation of the solution is recommended.
    Carbon filters can be partially regenerated with steam, which removes hydrocarbons and other adsorbed contaminants. Regeneration or bed change out is recommended when foam tests on the inlet and outlet streams show no improvement. This indicates carbon bed saturation.
    Filters may be located on the lean or rich solution side. Filtration on the rich side seeks to remove particles that are more insoluble under the rich solution conditions. It also prevents solids accumulation in the hot environment of the stripper.
    However, proper design and operating procedures for personnel protection during filter maintenance is mandatory when H2S may be present.
    Filtration equipment should be used continuously beginning with the first day of plant operations. When starting up the plant, the full flow filter, even if temporary, may prove its worth by removing the scale and other solid particles and allowing much quicker and easier start-up of the plant.

    6.6.2 Flash Tank
    Rich solution leaving the contactor may pass through a flash tank. A flash tank is more important when treating high pressure gas. Gases entrained in the rich solution will be separated.
    In addition, the amount of absorbed gas will be decreased because of the lower operating pressure of the flash tank. Using a flash tank will:
    Reduce erosion in rich/lean exchangers.
    Minimize the hydrocarbon content in the acid gas.
    Reduce the vapor load on the stripper.
    Possibly allow the off-gas from the flash tank to be used as fuel (may require sweetening).
    When heavy hydrocarbons are present in the natural gas, the flash tank can also be used to skim off the heavy hydrocarbons that were absorbed by the solution. Residence times for flash tanks in amine service vary from 3-10 minutes depending on separation requirements. Inlet gas streams containing only methane and ethane require shorter residence times. Rich gas streams require longer time for the dissociation of gas from solution and/or the separation of liquid phases.

    6.6.3 Corrosion
    Corrosion is an operating concern in nearly all sweetening installations. The combination of H2S and CO2 with water practically ensures that corrosive conditions will exist in portions of the plant. In general, gas streams with high H2S to CO2 ratios are less corrosive than those having low H2S to CO2 ratios. H2S concentrations in the ppmv range with CO2 concentrations of 2 percent or more tend to be particularly corrosive.
    Because the corrosion in sweetening plants tends to be chemical in nature, it is strongly a function of temperature and liquid velocity. The type of sweetening solution being used and the concentration of that solution has a strong impact on the corrosion rate. Increased corrosion can be expected with stronger solutions and higher gas loadings.
    Hydrogen sulfide dissociates in water to form a weak acid. The acid attacks iron and forms insoluble iron sulfide. The iron sulfide will adhere to the base metal and may provide some protection from further corrosion, but it can be eroded away easily, exposing fresh metal for further attack. High liquid velocities can erode the protective iron sulfide film with resulting high corrosion rates.
    CO2 in the presence of free water will form carbonic acid. The carbonic acid will attack iron to form a soluble iron bicarbonate which, upon heating, will release CO2 and an insoluble iron carbonate or hydrolize to iron oxide.
    In general, design velocities in rich solution piping should be 50% of those that would be used in sweet service. Because of the temperature relationship to corrosion, the reboiler, the rich side of the amine-amine exchanger, tend to experience high corrosion rates. Because of the low pH the stripper overhead condensing loop also tends to experience high corrosion rates.
    Acid degradation products also contribute to corrosion. A suggested mechanism for corrosion is that degradation products act as chelating agents for iron when hot. When cooled, the iron chelates become unstable, releasing the iron to form iron sulfide in the presence of H2S. Primary amines are thought to be more corrosive than secondary amines because the degradation products of the primary amines act as stronger chelating agents.
    Several forms of stress corrosion *****ing are possible in amine sweetening systems. Amine stress corrosion *****ing can occur and is worst in hot solutions, but *****ing can occur in cooler lines and both rich and lean streams. Wet sulfide *****ing and blistering can occur due to hydrogen generated in corrosion reactions. The hydrogen can collect at small inclusions in the steel which delaminate and then link in a step-wise pattern to create blisters. This is called HIC or hydrogen induced *****ing. Sometimes stress influences the *****ing to cause SOHIC or stress oriented hydrogen induced *****ing. HIC resistant steels are available. Seamless pipe is less prone to HIC than plate steels.
    Corrosion in alkaline salt processes, such as the hot carbonate process, has been reported to range from none to severe. Corrosion can be expected where CO2 and steam are released through flashing. Severe erosion can take place when carbonate solution strengths exceed 40% because of the tendency to form bicarbonate crystals when the solution cools.
    Many corrosion problems may be solved using corrosion inhibitors in combination with operating practices which reduce corrosion.
    Following are some guidelines to minimize corrosion.
    Maintain the lowest possible reboiler temperature.
    If available, use low temperature heat medium rather than a high temperature heat medium or direct firing. When a high temperature heat medium or direct firing for the reboiler is used, caution should be taken to add only enough heat for stripping the solution.
    Minimize solids and degradation products in the system through reclaimer operation and effective filtration.
    Keep oxygen out of the system by providing a gas blanket on all storage tanks and maintain a positive pressure on the suction of all pumps.
    Ensure deionized water or oxygen/chemical-free boiler condensate is used for make up water. If available, steam can be used to replace water loss.
    Limit solution strengths to minimum levels required for treating.
    Pipe solution exchangers for upflow operation with the rich solution on the tube side.
    Monitor corrosion rates with coupons or suitable corrosion probes.
    Maintain adequate solution level above reboiler tube bundles and fire tubes; a minimum tube submergence of 12" is recommended.

    Corrosion inhibitors used include high molecular weight amines and heavy metal salts. These inhibitors offer potential savings in both capital and operating costs for these special cases.

    6.6.4 Foaming
    A sudden increase in differential pressure across a contactor or a sudden liquid level variation at the bottom of the contactor often indicates severe foaming. When foaming occurs, there is poor contact between the gas and the chemical solution. The result is reduced treating capacity and sweetening efficiency, possibly to the point that outlet specification cannot be met.
    Some reasons for foaming are:
    Suspended solids
    Organic acids
    Corrosion inhibitors
    Condensed hydrocarbons
    Soap-based valve greases
    Makeup water impurities
    Degradation products
    Lube Oil
    Foaming problems can usually be traced to plant operational problems. Contaminants from upstream operations can be minimized through adequate inlet separation. Condensation of hydrocarbons in the contactor can usually be avoided by maintaining the lean solution temperature at least 10F above the hydrocarbon dew point temperature of the outlet gas.
    Temporary upsets can be controlled by the addition of antifoam chemicals. These antifoams are usually of the silicone or long-chain alcohol type.
    The following test for foaming should be run with the various types of inhibitors being considered for a given application. This test should give the operator an indication of which antifoam will be the most effective for the particular case. Place several drops of antifoam in 200 ml of treating solution contained in a 1000 ml cylinder. Bubble oil-free air through the solution at a constant rate. After five minutes have elapsed shut off the air and start a timer. Note the height of foam at the time the air was shut off and the amount of time required for the foam to break. The foam height is the difference between the height of the foam and the initial height of the liquid. The time for the foam to break is an indication of the stability of the foam. A comparison of antifoams will let the operator select which inhibitor will best solve his foaming problems. Between antifoam tests, care should be taken to clean the test cylinder thoroughly because a very small amount of inhibitor may affect the test.

    6.6.5 Materials
    Treating plants normally use carbon steel as the principal material of construction. Vessels and piping should be stress relieved in order to minimize stress corrosion along weld seams. Corrosion allowance for equipment ranges from 1/16" to 1/4", typically 1/8". In some instances, when corrosion is known to be a problem, or high solution loadings are required, stainless steel or clad stainless steel may be used in the following critical areas:
    1. Reflux condenser
    2. Reboiler tube bundle
    3. Rich/lean exchanger tubes
    4. Bubbling area of the contactor and/or stripper trays.
    5. Rich solution piping from the rich/lean exchanger to the stripper.
    6. Bottom 5 trays of the contactor and top 5 trays of stripper, if not all.
    Usually 304, 316, or 410 stainless steel will be used in these areas, even through corrosion has been experienced with 410 stainless in DEA service for CO2 removal in the absence of H2S.
    L grades are recommended if the alloys are to be welded.
    Controlling oxygen content to less than 0.2 ppmw is effective in preventing chloride SCC in waters with up to 1000 ppmw chloride content, at temperatures up to 570F. There has been an increased use of duplex stainless steels, and they have been successfully used in the water treatment industry to prevent chloride SCC in high chloride waters. This suggests duplex stainless steels could be utilized in amine plant service where high chloride content is expected. As with any specialty steel, proper fabrication techniques and welding procedures are required.

    6.7 Solid Bed Processes
    6.7.1 General Process Description
    A fixed bed of solid particles can be used to remove acid gases either through chemical reactions or through ionic bonding.
    This process flows the gas stream through a fixed bed of solid particles, which removes the acid gases and holds them in the bed. When the bed is spent, the vessel must be removed from service and the bed regenerated or replaced. Since the bed must be removed from service to be regenerated, some spare capacity is normally provided.

    Commonly used processes under this category are as follows:
    Acid gas removal through chemical reactions include:
    Iron Sponge Process, SulfaTreat, Zinc Oxide Process, Chemsweet, & PuraSpec
    Acid gas removal through Ionic bonding Adsorption include:
    Molecular Sieve Process, & impregnated activated carbon.

    6.7.2 Iron Sponge Process
    The iron sponge process is economically applied to gases containing small amounts of H2S (<300 ppm) operating at low to moderate pressures in the range of 25-500 psig. This process does not remove CO2.
    The reaction of iron oxide and H2S produces iron sulfide and water as follows:
    Fe2O3 +3H2S → Fe2S3 +3H2O Eq. 6-12
    FeO+H2S → FeS+H2O Eq. 6-13
    The reaction requires the presence of slightly alkaline water (pH 8-10) and a temperature below 110 0F (47 0C). If the gas does not contain sufficient water vapor, water may need to be injected into the inlet gas stream. The pH level can be maintained through the injection of caustic soda, soda ash, lime, or ammonia with the water.

    Although the presence of free alkalines enhances H2S removal, it also creates potential safety hazards and promotes the formation of undesirable salts, adding to capital costs.
    Ferric oxide is impregnated on wood chips,which produce a solid bed with a large ferric oxide surface area. Several grades of treated wood chips are available, based on iron oxide content. Ferric oxide wood chips are available in 6.5, 9.0,15.0, and 20 lbs iron oxide/bushel.Chips are contained in a vessel, and sour gas flows downward through the bed and reacts with the ferric oxide. Figure 6-11 shows a vertical vessel used in the iron sponge process.

    The bed can be regenerated with air; however, only about 60% of the previous bed life can be expected. The bed life of the batch process is dependent upon the quantity of H2S, the amount of iron oxide in the bed, residence time, pH, moisture content, and temperature.
    Ferric sulfide can be oxidized with air to produce sulfur and regenerate the ferric oxide. Regeneration must be performed with care because the reaction with oxygen is exothermic (i.e., gives off heat). Air must be introduced slowly so that the heat of reaction can be dissipated.
    If air is introduced quickly, the heat of reaction may ignite the bed.
    For this reason, spent wood chips should be kept moist when removed from the vessel. Otherwise, the reaction with oxygen in the air may ignite the chips and cause them to smolder.

    The reactions for oxygen regeneration are as follows:
    2Fe2S3 + 3O2 + 2H2O → 2Fe2O3.(H2O) + 6S+Heat Eq. 6-14
    4FeS +3O2 + 2xH2O → 2Fe2O3(H2O)x +4S +Heat Eq. 6-15
    S2 +2O2 → 2SO2 Eq. 6-16

    Some of the elemental sulfur produced in the regeneration step remains in the bed. After several cycles, this sulfur will cake over the ferric oxide, decreasing the reactivity of the bed and causing excessive gas pressure drop.
    Typically after 10 cycles, the bed must be removed from the vessel and replaced with a new bed.
    It is possible to operate an iron sponge with continuous regeneration by the introduction of small amounts of air in the sour gas feed. The oxygen in the air regenerates the iron sulfide and produces elemental sulfur. Although continuous regeneration decreases the amount of operating labor, it is not as effective as batch regeneration, and it may create an explosive mixture of air and natural gas. Due to the added costs associated with an air compressor, continuous regeneration generally does not prove to be the economic choice for the typically small quantities of gas involved.
    Cooler operating temperatures of the natural gas, for example, during the winter, create the potential for hydrate formation in the iron sponge bed.
    Hydrates can cause high-pressure drop, bed compaction, and flow channeling.
    When the potential for hydrates exists, methanol can be injected to inhibit their formation. If insufficient water is present to absorb the methanol, it may coat the bed, forming undesirable salts. Hydrocarbon liquids in the gas tend to accumulate on the iron sponge media, thus inhibiting the reactions.
    The use of a gas scrubber upstream of the iron sponge at a gas temperature slightly less than that of the sponge media may prevent significant quantities of liquids from condensing and fouling the bed.
    There has been a recent revival in the use of iron sponges to sweeten light hydrocarbon liquids. Sour liquids flow through the bed and are contacted with the iron sponge media and the reaction proceeds as above.

    Fig. 6-11 Iron oxide acid-gas treating unit.

    6.7.3 SulfaTreat
    The Sulfa-Treat process similar to iron sponge process. It is economically applied to gases containing small amounts of H2S. This process utilizes a proprietary iron oxide co-product mixed with inert powder to form a porous bed. Sour gas flows through the bed and reacts with iron as in iron-sponge reaction.
    The material of SulfaTreat is a dry, free-flowing granular substance used for selective removal of H2S and mercaptans from natural gas in the presence of CO2. It is not affected by CO2, and it does not produce sulfur or nitric oxide (NOx). Also SulfaTreat will not ignite or "cement up" in the vessel. Other advantages include longer bed life and lower cost.
    SulfaTreats particle size varies from 4 to 30 mesh and has a bulk density of 70 lb/ft3. These physical properties give uniform porosity and permeability, which offers small resistance to flow and resistance to bed compression at normal velocities.
    Applications for SulfaTreat include: natural gas treating, amine treater off gas, high concentration CO2 streams, and any other H2S-containing system.
    The reaction works better with saturated gas and at elevated temperature up to 1300F (54.40C). There is no minimum moisture or pH level required. The amount of bed volume required increases as the velocity increases and as the bed height decreases. Operation of the system below 400F (4.40C) is not recommended. Beds are not regenerated and must be replaced when the bed is spent.

    6.7.4 Zinc Oxide Process
    The equipment used in the zinc oxide process is similar to the iron sponge process. It uses a solid bed of granular zinc oxide to react with the H2S to form zinc sulfide and water as shown below
    ZnO+H2S → ZnS+H2O Eq. 6-17
    The rate of reaction is controlled by the diffusion process, as the sulfide ion must first diffuse to the surface of the zinc oxide to react. Temperatures above 2500F (1200C) increase the diffusion rate, which promotes the reaction rate. The strong dependence on diffusion means that other variables, such as pressure and gas velocity, have little effect on the reaction.
    The zinc oxide is contained in long thin beds to lessen the chances of channeling. The pressure drop through the beds is low. Bed life is a function of gas H2S content and can vary from 6 months to more than 10 years. Beds are often used in series to increase the level of saturation prior to change out of the catalyst. A spent bed is discharged by gravity flow through the bottom of the vessel.
    The zinc oxide process is seldom used due to disposal problems with the spent bed, which is classified as a heavy metal salt.

    6.7.5 Chemsweet
    Chemsweet is the name for another batch process for the removal of H2S from natural gas. Chemicals used are a mixture of zinc oxide, zinc acetate, water, and a dispersant to keep solid particles in suspension. Natural gas is bubbled through the solution where H2S reacts with zinc oxide. Though several reactions take place in solution, the net result is that zinc oxide reacts with H2S to form zinc sulfide and water.
    Chemsweet can treat gas streams with H2S concentration up to 400 ppmv. and has been operated between pressures of 75 and 1,400 psig. Mercaptan concentrations in excess of 10% of the H2S concentration in the gas stream can cause a problem. Some of the mercaptans will react with the zinc oxide and be removed from the gas. The resulting zinc mercaptides [Zn(OH)RH] will form a sludge and possibly cause foaming problems.

    6.7.6 PuraSpec
    Johnson Matthey Catalysts supplies the PuraSpec range of processes and products for desulfurization of hydrocarbon gases and liquids. The processes use fixed beds of granular, metal oxide-based chemical absorbents which are developments of the high temperature zinc oxide used for purification of hydrocarbon feedstocks to steam reformers in ammonia, hydrogen, and methanol plants. PuraSpec absorbents are effective at temperatures down to 32F, so no added heat is necessary, and are in service at pressures from atmospheric, treating vent gases, to 1800 psi treating dense phase gas feed to a gas processing plant.
    PuraSpec units are in service treating natural gas to pipeline or petrochemical specifications. Because the absorbents remove H2S and COS irreversibly, they are best suited to polishing duties.

    6.7.8 Molecular Sieve Process
    Acid gases, as well as water, can be effectively removed by physical adsorption on synthetic zeolites. Applications of acid gas removal using molecular sieve are limited because water displaces acid gases on the adsorbent bed. (Chapter 5 provides more details on adsorption and its use in dehydration.)
    Crystalline structure of the solids provides a very porous material having uniform pore size. Within the pores the crystalline structure creates a large number of localized polar charges called active sites. Polar gas molecules, such as H2S and water vapor, which enter the pores, form weak ionic bonds at the active sites. Nonpolar molecules, such as paraffin hydrocarbons, will not bond to the active sites.
    Carbon dioxide molecules are about the same size as H2S molecules, but are nonpolar. CO2 will enter the pores but will not bond to the active sites. Small quantities of CO2 will be removed by becoming trapped in the pores by bonded H2S or H2O molecules blocking the pores.
    CO2 will obstruct the access of H2S and H2O to the active sites, thus decreasing the overall effectiveness of the molecular sieve. Beds must be sized to remove all H2O and provide for interference from other molecules in order to remove all H2S.
    The adsorption process usually occurs at moderate pressure. Ionic bonds tend to achieve an optimum performance near 450 psig but can operate in a wide range of pressures.

    Hydrogen sulfide can be selectively removed to meet 0.25 grain/100 scf (4ppm) specification. However, this reduction requires regeneration of the bed at 600F (315C) for extended time (Usually one hour or more depending upon process conditions) with the potential for COS formation if 4A is used (Molecular sieves are classified by their nominal pore diameter in Angstroms. See Chapter 5 for details.).
    The problem of COS formation during processing according to the reaction in Equation 6-18, has been extensively studied.
    H2S + CO2 ↔ COS + H2O Eq. 6-18
    Molecular sieve products have been developed that do not catalyze COS formation. The central zone in the regeneration cycle is most favorable to COS formation.
    Regeneration of a molecular sieve bed concentrates the H2S into a small regeneration stream which must be treated or disposed of.

    Figure 6-12 shows a typical flow diagram for removal of H2S from natural gas. The configuration is similar to that for dehydration but with the significant difference that the regeneration gas contains high quantities of H2S as well as water as it leaves the adsorbent bed and, thus, must be treated. (Chapter 5 presents a thorough discussion of adsorption.) The flow configuration shows the first bed in the adsorption cycle, the second bed cooling down after regeneration, and the third bed undergoing regeneration with hot gas.

    Fig. 6-12. Integrated Natural Gas Desulfurization Plant
    Chi and Lee (1973) studied the coadsorption of H2S, CO2, and H2O on a 5A molecular sieve from a natural gas mixture under a variety of conditions.
    Figure 6-13, from their paper, shows a typical concentration versus time curve.
    In the figure, y is the concentration in the exit stream and yo is the concentration in the inlet to the bed. The gas that entered the bed was saturated with H2O and contained both CO2 (1.14 mol%) and H2S (0.073 mol%.). Because the CO2 content of the gas was 15.6 times that of the H2S, the bed quickly saturated with CO2, and its breakthrough was almost instantaneous. As the H2S was adsorbed and moved down the column, it displaced the CO2 and, consequently, after approximately 30 minutes, the CO2 exit concentration peaked at a value greater than its inlet concentration.
    The same phenomenon occurs when the H2S is displaced by the water.
    Because the H2S must be removed to extremely low levels, the bed is effectively exhausted from the perspective of H2S purification shortly after H2S breakthrough occurs and, thus, well before the bed is totally saturated.

    Fig. 6.13 Effluent H2S and CO2 concentration from adsorption bed as a function of Time.

    In general, the sieve bed can be designed to dehydrate and sweeten simultaneously.
    The Engineering Data Book (GPSA) notes that a key point in adsorber design is to properly design for treatment of the regeneration gas because the peak H2S concentration may be 30 times the H2S concentration in the feed.
    Care should be taken to minimize mechanical damage to the solid crystals as this will decrease the beds effectiveness. The main cause of mechanical degradation is the sudden pressure and/or temperature changes that may occur when switching from adsorption to regeneration cycles. Proper instrumentation can significantly extend bed life.
    The molecular sieve process is limited to small gas streams operating at moderate pressures. It is generally used for polishing applications following one of the other processes.

    6.7.9 Oxorbon
    Oxorbon is an alternative solid bed material which consists of activated carbon impregnated with potassium iodide (KI). Donau Carbon markets such a carbon for the removal of H2S and mercaptans. The adsorbed H2S is converted to elemental sulfur by catalytic reaction under the presence of oxygen. The resulting sulfur is fixed on the pores of the activated carbon. Sulfur loadings as high as 60% of the carbon mass have been reported and sulfide concentrations below 1 ppmv are claimed.

    6.8 Direct Conversion Processes (Liquid Redox)
    Chemical and physical solvent processes remove acid gas from the natural gas stream but release H2S and CO2 when the solvent is regenerated. The release of H2S to the atmosphere is limited by environmental regulations. Acid gases could be routed to an incinerator/flare, which would convert the H2S to SO2. Environmental regulations restrict the amount of SO2 vented or flared. Direct conversion processes use chemical reactions to oxidize H2S and produce elemental sulfur. These processes are generally based either on the reaction of H2S and O2 or H2S and SO2. Both reactions yield water and elemental sulfur. These processes are licensed and involve specialized catalysts and/or solvents.
    The redox agent is then regenerated by reaction with air in an oxidizer vessel.
    Liquid oxidation−reduction (redox) processes use iron as an oxidizing agent in solution. The process involves four basic steps:
    1. Removal of H2S from the gas by absorption into a caustic solution
    2. Oxidation of the HS− ion to elemental sulfur via the oxidizing agent
    3. Separation and removal of sulfur from the solution
    4. Regeneration (i.e., oxidation) of the oxidizing agent by use of air

    The processes share the general chemistry shown below.
    H2S absorption using an alkaline solution:
    H2S (g) ↔ H2S (soln) Eq 6-19
    H2S (soln) ↔ H+ + HS− Eq 6-20

    Note that this is a non-selective, basic scrubbing solution; thus, additional acidic components (CO2 , HCN) of the sour gas will be soluble in the scrubbing solution. While CO2 is partially absorbed in the alkaline solution, it does not participate in the redox reactions. CO2 may however increase the rate of consumption of caustic or buffering compounds. The fate of other reduced sulfur species (COS, CS2, RSH, RSR) depends on the particular process being considered.
    Conversion to Elemental Sulfur:
    HS− + H+ + 1/2 O2 (soln) → H2O (l) + S Eq 6-21

    Note that in the absence of an auxiliary redox reagent (ARR) the reaction shown as Eq 21-24 is slow and nonspecific. The addition of an (ARR) increases the rate of reaction and directs the oxidation to elemental sulfur.
    HS− + ARR(OX) → S + H+ + ARR (RED) Eq 6-22

    where (OX) denotes the oxidized form and (RED) denotes the reduced form of the ARR.
    Regeneration of the Spent ARR Using Air:
    O2 (g) ↔ O2 (soln) Eq 6-23
    ARR (RED) + 1/2 O2 (soln) + 2 H+ → ARR(OX) + H2O(l) Eq 6-24

    The overall simplified chemistry of liquid redox processes is thus:
    H2S(g) + 1/2 O2 (g) → H2O(l) + S Eq 6-25

    While the oxidizing agent forms the desired reaction, it also reacts with the sulfur species to form metal sulfides that precipitate from solution. To avoid the metal sulfide reaction, the solution contains chelating agents (organic compounds that bind with the metal ion to restrict its reactivity but still permit electron transfer for oxidation−reduction reactions). A common chelating agent for iron is EDTA (ethylenediaminetetraacetic acid). Several processes, including Lo-Cat II, SulFerox , and Sulfint-HP, use iron and chelating agents.
    6.8.1 Stretford Process
    The Stretford process involves the use of vanadium salts as the ARR. The process has been extensively used in Europe. However environmental concerns around the discharge of vanadium compounds has limited its use.
    Figure 6.14 shows a simplified diagram of the Stretford process. The gas stream is washed with an aqueous solution of sodium carbonate, sodium vanadate, and anthraquinone disulfonic. An oxidized solution is delivered from the pumping tank to the top of the absorber tower where it contacts the gas stream in a counter-current flow.
    The bottom of the absorber tower consists of a reaction tank from which the reduced solution passes to the solution flash drum, which is situated above the oxidizer. The reduced solution passes from here into the base of the oxidizer vessel. Hydrocarbon gases, which have been dissolved in the solution at the plant pressure, are released from the top of the flash drum.
    Air is blown into the oxidizer, and the main body of the solution, now reoxidized, passes into the pumping tank. The sulfur is carried to the top of the oxidizer by froth created by the aeration of the solution and passes into the thickener.
    The function of the thickener is to increase the weight percent of sulfur that is pumped to one of the alternate sulfur recovery methods of filtration, filtration and autoclaves, centrifugation or centrifugation with heating.
    Chemical reactions involved are:
    H2S+Na2CO3 → NaHS+NaHCO3 Eq 6-26
    Sodium carbonate provides the alkaline solution for initial adsorption of H2S and the formation of hydrosulfide (HS). The hydrosulfide is reduced in a reaction with sodium meta vanadate to precipitate sulfur
    HS- +V+5 → S +V+4 Eq 6-27
    Anthraquinone disulfonic acid (ADA) reacts with 4-valent vanadium and converts it back to 5 valent
    V+4 +ADA → V+5 +ADA (reduced) Eq 6-28
    Oxygen from the air converts the reduced ADA back to the oxidized state as shown below:
    Reduced ADA+O2 → ADA+H2O Eq 6-29
    The overall reaction is
    2H2S+O2 → 2H2O+2S Eq 6-30

    6.8.2 Lo-Cat process
    This is a liquid-redox process. The LO-CAT process is a patented, wet scrubbing, liquid redox system that uses a chelated iron solution to convert H2S to innocuous, elemental sulfur. It does not use any toxic chemicals and does not produce any hazardous waste byproducts. In general, the LO-CAT process can directly treat a gas stream, or treat the H2S containing stream from an Acid Gas Removal unit, although the direct treat capability is limited to low pressure streams.
    Processes employ high iron concentration reduction-oxidation technology for the selective removal of H2S (not reactive to CO2) to <4 ppm in both low and high pressure gas streams. Figure 6-15 shows a simplified flow schematic of the LO-CAT process.
    Acid gas stream is contacted with the solution where H2S reacts with, and reduces the chelated iron and produces elemental sulfur. The iron is then regenerated by bubbling air through the solution. Heat is not required for regeneration. The reactions involved are exothermic (give off heat):

    Absorption/reduction: 2Fe3+ +H2S → 2Fe2+ +S+2H+ Eq 6-31
    Regeneration/oxidation: 2Fe2+ + O2 +2H+ → 2Fe3+ +H2O Eq 6-32
    Overall chemistry: H2S + O2 → S + H2O Eq 6-33

    Fig 6.14. Simplified flow schematic of the Stretford process.

    Fig. 6.15. Simplified flow schematic of the LO-CAT process.

    6.8.3 Sulferox process
    SulFerox is a Shell proprietary Iron Redox process, whereby a sour gas stream, containing hydrogen sulphide, is contacted with a liquid, containing soluble ferric (Fe3+) ions. In the process the H2S is oxidized to elemental sulphur and the Fe3+ is reduced to ferrous (Fe2+) ions. The system is regenerable, Fe2+ is subsequently reconverted to Fe3+ by oxidation with air. Sulphur is recovered from the aqueous solution as a moist cake.
    The process offers low capital and operating costs through the use of a high concentration iron chelate solution and effective control of chelate degradation. The process also offers a patented contactor design to improve the overall efficiency of the process which further reduces capital costs. The optimum application for SulFerox is in the one to twenty tons per day range of recovered sulfur.
    Typical Sulferox process equipment is illustrated in Figure 6.16. In this process, the sour gas is contacted in a small contactor in co-current flow with a water solution containing about 4% ferric iron ions, held in solution by proprietary chelate ligand. The H2S ionizes in the solution, and the ferric ions exchange electrons with the sulfur ions to form ferrous ions and elemental sulfur. The gas and the solution leave the contactor and are flowed into a separator. The gas out of the separator is sweet and requires further treating for dewpoint control. The solution is flowed into additional vessels for separating the elemental sulfur and for restoring the ferrous iron ions back to the active ferric iron state by contacting the solution with air.
    In general, the Sulferox process can directly treat a gas stream, or treat the H2S containing stream from an Acid Gas Removal unit, although the direct treat capability is limited to low pressure streams.

    Fig. 6.16. Simplified flow schematic of the sulferox process.

  7. Re: Basics of Gas Field Processing Book "Full text"

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

    Fig. 6-21. Claus Process Variations

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

    6.11.6 Membrane Design Considerations
    Process Variables Affecting Design are: Flow Rate
    A maximum acceptable feed gas rate per unit area applies to the membrane, and required membrane area is directly proportional to the flow rate. Membrane units perform well at reduced feed rates, but their performance drops when design flow rates are exceeded. Additional modules are added in parallel to accept higher flow rates.
    The percentage of hydrocarbon losses (hydrocarbon losses/feed hydrocarbons) remains the same at different flow rates. Operating Temperature
    Increased operating temperature increases permeability but decreases selectivity.
    Membrane area requirement is decreased, but the hydrocarbon losses and recycle compressor power for multistage systems are increased (Figure 6-29).
    Because membranes are organic polymers, they have a maximum operating temperature that depends upon the polymer used. Exceeding this temperature will degrade membrane material and shorten the useful life of the unit. Feed Pressure
    An increase in feed pressure decreases both membrane permeability and selectivity, but at the same time creates a greater driving force across the membrane that results in a net increase in permeation through the membrane and a decrease in the membrane area requirements (Figure 6-30). Increasing the maximum operating pressure results in a less expensive and smaller system. Limiting factors are the maximum pressure limit for the membrane elements and the cost and weight of equipment at the higher pressure rating.

    Fig. 6-29. Effect of operating temperature.

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

    Fig. 6-31. Effect of permeate pressure.

    Desirable to achieve the lowest possible permeate pressure Important consideration when deciding how to further process the permeate stream.
    For example, if permeate stream must be flared, then the flare design must be optimized for low pressure drop. If permeate stream must be compressed to feed the second membrane stage or injected into a well, the increased compressor horsepower and size at lower permeate pressure must be balanced against the reduced membrane area requirements. CO2 Removal
    For a constant sales gas CO2 specification, an increase in feed CO2 increases membrane area requirement and increases hydrocarbon losses (more CO2 must permeate, and so more hydrocarbons permeate). This is shown in Figure 6-32.

    Fig. 6-32. Effect of CO2 removal.

    Membrane area requirement is determined by the percentage of CO2 removal rather than the feed or sales gas CO2 specifications themselves.
    For example, a system for reducing a feed CO2 content from 10% to 5% is similar in size to one reducing the feed from 50% to 25%, or one reducing a feed from 1% to 0.5%, if all have a CO2 removal requirement of about 50%.
    Traditional solvent or absorbent-based CO2 technologies have the opposite limitation.
    Their size is driven by the absolute amount of CO2 that must be removed. For example, a system for CO2 removal from 50% to 25% is substantially larger than one reducing CO2 from 1% to 0.5%. For this reason, using membranes for bulk CO2 removal and using traditional technologies for meeting low CO2 specifications makes a lot of sense. Depending on the application, either one or both of the technologies could be used.
    An increase in CO2 content in feed gas of an existing membrane plant will results in sales gas with higher CO2 content. An additional membrane area can be installed to meet the sales gas CO2 content, although with increased hydrocarbon losses. For example, if heater capacity is available, the membranes can be operated at a higher temperature to also increase capacity.
    If an existing non-membrane system must be de-bottlenecked, installing a bulk CO2 removal system upstream of it makes good sense. Environmental Regulations
    Environmental regulations dictate what can be done with the permeate gas, specifically whether it can be vented (cold or hot vent) to the atmosphere or flared either directly or catalytically. Ninety five to ninety nine percent CO2 yields low Btu/scf content (flare requires a minimum of 250 Btu/scf to burn). Location
    Location often dictates a number of other issues, such as space and weight restrictions, level of automation, level of spares that should be available, and single versus multistage operation.
    Fuel requirements can be obtained upstream of the membrane system, downstream of the pretreatment system, downstream of the membrane, and from the recycle loop in multistage systems.

  8.    Spons.

  9. Re: Basics of Gas Field Processing Book "Full text"

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

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

    Table 6.14. Typical membrane feed and permeate gas analysis

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

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

    [link Point to another website Only the registered members can access] Fig. 6-34. A two stage CO2/CH4 membrane separation process.

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

    Fig 6.36 Schematic of membrane-pretreating equipment (Traditional Pretreatment). Traditional Pretreatment
    A common method for pretreating the feed gas to a membrane system is shown in Figure 6-36. The coalescing filter removes any entrained liquids; the adsorbent bed takes out trace contaminants such as volatile organic compounds (VOC); the particulate filter removes any dust from the adsorbent bed; and the heater superheats the gas to prevent liquid formation in the membrane unit. The system shown has the following disadvantages:
    The adsorbent bed is the only unit that removes heavy hydrocarbons. Consequently, if the gas contains more heavy hydrocarbons than anticipated, or in the event of a surge of these materials, the adsorbent bed may become saturated in a relatively short time, and thus allow heavy hydrocarbons to contact the membrane.
    Only the heater provides superheat, and, consequently, if this unit fails, the entire membrane system must be shut down.
    Other pretreatment methods that address the disadvantages discussed above.
    A chiller may be included to reduce the dew point of the gas and the heavy hydrocarbon content. Because chilling does not completely remove all heavy hydrocarbons, an adsorbent guard bed is still required. If deep chilling is necessary, steps must be taken to prevent hydrates from forming, either by dehydrating the gas upstream or by adding hydrate inhibitors. If inhibitors are added, they may need to be removed downstream of the chiller because some inhibitors may damage the membrane.
    The turbo-expander serves the same purpose as a chiller, but has the benefit of being a dry system. It is smaller and lighter than the refrigeration system. A disadvantage is the net pressure loss, which must be taken up by the export compressor.
    Glycol Unit
    The glycol unit is added upstream of the chiller to prevent hydrate formation or freeze-up. An adsorbent guard bed is still required to remove heavy hydrocarbons but must be larger than it would normally be because it must also remove glycol carried over from the adsorber vessels. Enhanced Pretreatment
    It is common for an initial design, based on an extended gas analysis, to differ from actual analysis after the membrane system has been started up. For example, feed gas may be heavier than originally anticipated. Pretreatment systems may not have sufficient flexibility to handle a wide departure from design. Adsorbent beds may become fully saturated within a short time, leading to performance degradation. Preheaters may not be large enough to achieve feed temperatures that are much higher than designed.
    A standard way to handle a gas that is heavier than expected is to operate the membranes at a higher temperature. Temperature increase increases the margin between gas dew point and operating temperature and thus prevents condensation in the membrane.
    Figure 6.37 shows an enhanced pretreatment scheme that is more suitable for cases where one or more of the following is expected:
    A wide variation in feed gas content
    A significant amount of heavy hydrocarbons or other contaminants; or
    A feed gas that may be heavier than analyzed, based on the known information from nearby wells or other locations.

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

    Fig. 6.37. Enhanced pretreatment flow scheme.

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

    6.11.9 Membrane Advantages & Disadvantages
    Recently, several very large membrane systems including some treating in excess of 300MMscfd of gas containing over 30% CO2 have been successfully implemented as a result of advances in pretreatment technology. Membrane system suppliers include Grace, Kvaerner Process Systems, Natco-Cynara,UOP, and Ube.
    Natural gas sweetening and dehydration using membranes often offers significant advantages over the more conventional methods such as amine treating, physical solvents, and solid adsorbents.
    Membranes are particularly attractive when the pressure of the feed gas is high (over 500 psig) and/or the CO2 content of the gas to be treated is high (over 10%). Advantages
    Lower Capital Cost Capital Expenditure (CAPEX)
    Membrane systems are skid-mounted, except for larger pretreatment vessels.
    Scope, cost, and time required for site preparation are minimal. Installation costs are significantly lower than alternative technologies, especially for remote areas and offshore installations.
    Membrane systems do not require the additional facilities, such as solvent storage and water treatment needed by other processes.

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

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

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

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

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

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

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

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

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

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

    Ideal for Remote Locations
    Many of the factors mentioned above make membrane systems a highly desirable technology for remote locations where spare parts are rare and labor unskilled. Solvent storage and trucking, water supply, power generation (unless a multistage system is installed), or extensive infrastructure are not required. Disadvantages
    Clean feed: Pretreatment of the feed to the membrane to remove particulates and liquids is generally required
    Gas compression: Because pressure difference is the driving force for membrane separation, considerable recompression may be required for either or both the residue and permeate streams
    Generally higher hydrocarbon losses than solvent systems.
    Membrane materials are not suitable for high hydrogen sulfide partial pressures and applications for bulk H2S removal is not practical. According to Kvaerner Process Systems the maximum permissible H2S partial pressure is around 20 psia for present membrane materials.

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

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

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

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

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

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

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

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

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

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

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

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

  10. Re: Basics of Gas Field Processing Book "Full text"

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

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

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

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

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

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

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

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

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

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

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

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

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

    Fig. 6-40. Process selection chart for CO2 removal with no H2S present.

    [link Point to another website Only the registered members can access] Fig. 6-41.Process selection chart for H2S removal with no CO2 present.

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

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

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

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

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

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

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

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

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

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

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

    Step 5
    The amount of iron oxide, which is impregnated on the wood chips, is normally specified in units of pounds of iron oxide (Fe2O3) per bushel. Common grades are 6.5, 9, 15, or 20 lbs Fe2O3/bushel. Theoretical bed life for the iron sponge between replacements is determined from
    tc = 3.14 x 10-8 Fe d2 H E/ (Qg x H2S) Eq. 6-45
    where tc = cycle time, days; Fe = iron sponge content, lbs Fe2O3/bushel; E = efficiency (0.65-0.8).

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

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

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

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

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

    LDEA = 192 Qg XA / c ρ AL Eq. 6-47
    LMEA = MEA solution circulation rate, gpm; - LDEA = DEA solution circulation rate, gpm;
    Qg = gas flow rate, MMscfd); - XA = required reduction in total acid gas fraction, moles acid gas removed/mole inlet gas.
    Note: XA represents moles of all acid components, that is, CO2, H2S, and meracaptans, as MEA and DEA are not selective;
    c = amine weight fraction, kg amine/kg solution (lbs amine/lbs solution);
    ρ = solution density, lbs/gal; - AL = acid gas loading, mole acid gas/mole amine.
    The specific gravity of amine solutions at various amine concentrations is in Figure 6-44.
    For design purposes, the following solution strengths and loadings are recommended to provide an effective system without excessive corrosion rates:
    MEA solution strength20 wt% MEA
    DEA solution strength35 wt% DEA
    MEA acid gas loading0.33 mol acid gas/mol MEA
    DEA acid gas loading0.5 mol acid gas/mol DEA
    Density of MEA8.41 lbs/gal
    Density of DEA8.71 lbs/gal
    Using the recommended concentrations and specific gravities at 20 0C from Figure 6-44:
    20%MEA = 1.008SG = 1.008 x 8.34 lbs/gal
    = 8.41 lbs/gal = 8.41 x 0.20
    = 1.68 lbs MEA/gal
    =1.68/61.08 = 0.028 mol MEA/gal
    35%DEA = 1.044SG = 1.044 x 8.34 lbs/gal
    =8.71 lbs/gal = 8.71 lbs/gal x 0.35
    = 3.05 lbs DEA/gal
    = 3.05/105.14 = 0.029 mol DEA=gal
    Using these design limits, the circulation rates required can be determined from Equations (6-48) and (6-49):
    LMEA = 202 Qg XA Eq. 6-48

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

    Amine Reboiler
    The reboiler duty can be estimated as follows:
    qerb = 72,000 LMEA Eq. 6-50
    qerb = 60,000 LDEA Eq. 6-51
    qreb = reboiler duty, (Btu/h) - LMEA = MEA circulation rate, gpm - LDEA = DEA circulation rate, gpm.
    The operating temperature for amine reboilers is determined by the operating pressure and the lean solution concentration. Typical reboiler temperature ranges are as follows:
    MEA reboiler 225-260 0F (107-127 0C)
    DEA reboiler230-250 0F (110-121 0C)
    For design purposes, the reboiler temperature for a stripper operating at 10 psig (69 kPa) can be assumed to be 245 0F (118 0C) for 20% MEA, and 250 0F (121 0C) for 35% DEA. Boiling point versus solution concentration curves at various pressures are shown in Figure 6-45.

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

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

  11. Re: Basics of Gas Field Processing Book "Full text"

    Hydrocarbon Recovery - Chapter 7

    Fundamentals of Oil and Gas Processing Book
    Basics of Gas Field Processing Book
    Prediction and Inhibition of Gas Hydrates Book
    Basics of Corrosion in Oil and Gas Industry Book

    [link Point to another website Only the registered members can access]

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

    Chapter 7

    Hydrocarbon Recovery

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

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

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

    7.2.1 External Refrigeration
    External refrigeration is used to cool the gas stream to recover a significant amount of C3+ and to lower gas temperatures as the gas goes into other stages of hydrocarbon recovery. It may be the only source of refrigeration when inlet pressure is low. Adsorption and vapor compression refrigeration are used in special situations. However, vapor compression using propane as the working fluid is the most common in gas plants. (LNG facilities also use ethane or ethylene as well as hydrocarbon mixtures as refrigerants. Propane Refrigeration Process
    The refrigeration cycle consists of four steps that are depicted on the pressure− enthalpy chart in Figure 7.1:
    1. Compression of saturated refrigerant vapor at point A to a pressure well above its vapor pressure at ambient temperature at point B
    2. Condensation to point C by heat exchange with a cooling fluid, usually air.
    3. Expansion through a valve (Joule-Thomson expansion) to cool and condense the refrigerant to point D
    4. Heat exchange with the fluid to be cooled by evaporation of the refrigerant back to point A.

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

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

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

    Compression Step: Cycle analysis begins with propane vapor entering the compressor as a vapor at 14.5 psia (1 bar) and approximately −400F (−400C), where it is compressed to 250 psia (17 bar) (point A to point B in fig. 7.1.
    Condensation Step: The warm gas goes to an air- or water-cooled condenser, where the propane cools to 100 to 1200F (38 to 500C), totally condenses, and collects in a receiver (point B′ to point C in Figure 7.1).
    Expansion Step: Propane liquid leaves the receiver and flashes through a J-T valve, where the temperature and pressure drop to −400F (−400C) and 16 psia (1 bar) (point C to point D). No change occurs in the enthalpy, but the temperature drops to the saturation temperature of the liquid at the expansion-discharge pressure.
    Refrigeration Step: The cold propane then goes to a heat exchanger, the chiller, where it cools the process stream by evaporation (point D to point A in Figure 7.1).
    Because the propane in the chiller is evaporating, and a minimal heat exchange occurs between cold propane vapor and the inlet gas, the inlet and outlet propane temperature remains constant. The propane returns to the compressor suction slightly above −400F (−400C).
    In the past, most chillers were the kettle type in which propane is on the shell side and the liquid level is maintained above the tube bundle. Now, other high performance heat exchangers (e.g., plate-fin) are used. The chiller typically has two zones of heat transfer. The first is exchange of boiling propane with gas above its dew point and will involve only sensible heat. The second zone has condensing vapors from the process stream and boiling propane, which gives a much higher overall heat-transfer coefficient. To complete the cycle, the propane vapors leave the chiller and go to the suction drum before being compressed again. Alternate Process Configurations
    Thermodynamics dictate that to minimize the refrigeration work (i.e., compression required) heat from the chiller should be removed at as high a temperature as possible. One way to reduce compressor duty per unit of refrigeration duty is to multistage the refrigeration process by removal of process heat at more than one stage. Figure 7.3 shows a two-stage system, with representative operating conditions. In this system, the condensed propane stream expands to about 62 psia (4.3 bar) and 25F (−4C) in two parallel J-T valves. One of the expanded streams goes to a chiller before going to the suction surge drum for the second stage of the compressor.
    Vapor from the surge drum goes to the second stage of the compressor, while the liquid goes through a second J-T expansion to provide the low-temperature cooling. Pressures are adjusted so that the compression ratio is equal in both compressor stages. Table 7.1 shows the significant savings in going to two stages.
    Multistaging reduces work requirements by removing heat from the process stream at different temperatures. An alternative is removal of heat from the refrigerant before it is expanded. Refrigerant sub-cooling is sometimes used by exchange of the propane that leaves the receiver with a portion of the cold liquid propane.
    About 10% reduction in recirculation rate to the second stage of two-stage refrigeration is obtainable when the propane from the condenser is cooled 10F (5C).

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

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

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

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

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

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

    Fig. 7-4 Cutaway view of a turboexpander.

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

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

    7.2.3 Heat Exchange
    Most heat exchangers in a gas plant operating at or above ambient temperature are conventional shell and tube type and are ideal for steam and hot oil systems where fouling occurs. They are relatively inexpensive and easy to maintain because the tube bundle can be removed and tubes cleaned or replaced as needed.
    Where the fluids are clean and fouling does not occur, such as in gas−gas exchangers, compact heat exchangers are ideal. This section briefly discusses two kinds, brazed-aluminum plate-fin heat exchangers and printed circuit heat exchangers, which commonly are used in gas processing. Plate-Fin Exchangers
    Cryogenic facilities have made extensive use of brazed-aluminum plate-fin heat exchangers since the 1950s. Instead of a shell and tube configuration, these units consist of channels formed by a thin sheet of aluminum pressed into a corrugated pattern (the fin) sandwiched between two aluminum plates. Each layer resembles the end view of corrugated cardboard. The fin channels may be straight or may have a ruffled or louvered pattern to interrupt the straight flow path.
    Advantages of plate-fin exchangers include:
    Light weight.
    Excellent mechanical strength at subambient temperatures (used in liquid helium service [−4520F (−2680C)]). Can operate at pressures up to 1,400 psig (96 barg).
    High heat transfer surface area. Up to six times the surface area per unit volume of shell and tube exchanger and 25 times the area per unit mass.
    Complex flow configurations. Can handle more than 10 inlet streams with countercurrent, crossflow, and counter crossflow configurations.
    Close temperature approaches. Temperatures of 3F (1.7C) for single-phase fluids compared with 10 to 15F (6 to 9C) for shell and tube exchangers and 5F (2.8C) for two-phase systems.
    Drawbacks and limitations of the exchangers include:
    Single-unit construction. Repair can be more costly and time consuming than with shell and tube exchangers.
    Maximum operating temperature of approximately 150F (~85C), although special designs go to 400F (205C).
    Narrow channels. More susceptible to plugging, and fine mesh screens are needed where solids may enter. Components that might freeze out, water, CO2, benzene, and p-xylene, must be in sufficiently low concentrations to avoid plugging. The exchangers can be difficult to clean if plugging occurs.
    Less rugged. Does not accept rough handling or high pipe stress on nozzles.
    Limited to fluids noncorrosive to aluminum. Caustic chemicals are corrosive but not corroded by acid gases, unless free water is present.
    Susceptible to mercury contamination. Mercury amalgamates with aluminum to destroy mechanical strength.
    Susceptible to thermal shock. Maximum rate of temperature change is 4F/min (2C/min), and maximum difference between two streams is 55F (30C). Printed Circuit Heat Exchangers
    Another heat exchanger type, the printed circuit heat exchanger (PCHE) is used in clean service. This technology is relatively new, commercialized in the 1980s, but hundreds of units are in service. Like electronic printed circuits, heat transfer passages are etched in plates, and the plates are bonded together by diffusion bonding. Unlike the brazed-aluminum exchangers, they are rugged and, depending on materials of construction, go to high temperatures and pressures but can still handle complex flow schemes that involve many streams. Heat transfer passage sizes range from microchannels (less than 8 mil, 200 microns) to minichannels (0.12 in, 3 mm) to provide high heat transfer surface areas. Heat transfer area per unit volume can be 800 compared with 500 for plate-fin exchangers. Offshore operations employ PCHEs in many applications because they offer comparable heat transfer at comparable pressure drops at significantly less size and at one fifth the weight.

    7.2.4 Fractionation
    In addition to conventional distillation columns, two other types of distillation columns are commonly found in gas plants: stabilizers and demethanizers. Stabilizers are stripping columns used to remove light ends from NGL streams. Demethanizers are also stripping columns to remove methane from the NGL bottoms product. Demethanizers also act as the final cold separator, a collector of cold NGL liquids, and source of recovering some refrigeration by cooling warm inlet streams. Stabilizers
    The primary focus of dew pointing or fuel conditioning is to obtain a leaner gas. However, the by-product is a liquid phase that contains a substantial amount of volatiles. To make the liquid product easier to store and to recover more light ends for fuel or sales gas, many of the systems will stabilize the liquid by passing it through a stabilizer column. The stabilizer feed typically enters at the top of a packed or tray column and no reflux occurs. To increase stripping of light ends, the column pressure will be lower than that of the gas separator that feeds the column. In some cases, a stripping gas may be added near the bottom of the column in addition to the externally heated reboiler installed to provide additional vapor flow and enhance light-ends removal. This feature usually comes as an increased operating cost because the gas from the stripper is at low pressure and must be recompressed if put back into the inlet gas stream upstream of the gas treating unit. Demethanizer
    A distinguishing feature of gas plants with high ethane-recovery rates is the demethanizer. The column differs from usual distillation columns in the following ways:
    It has an increased diameter at the top to accommodate the predominately vapor feed to the top tray.
    It is typically primarily a stripping column, with no traditional condenserreflux stream.
    It may have several liquid feed inlets further down the column that come from low-temperature separators.
    It has a large temperature gradient; over 170F (75C) is common.
    The column serves two main functions: it acts as a flash drum for the top feed, which comes in as a cold, two-phase stream, and it removes methane from the bottoms product. Depending upon the plant configuration, the feed may be from a turboexpander, a J-T valve, or a heat exchanger. In some configurations the columns have reflux, but many demethanizers have no reflux.
    The NGL bottoms product is usually continuously monitored for methane content, which typically is kept below 0.5 liquid vol% of the ethane, on a C3+ free basis.
    The top of the column usually operates in the −175 to −165F (−115 to −110C) range, with pressures in the 200 to 400 psig (14 to 28 barg) range.

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

    7.3.1 Dew Point Control and Fuel Conditioning
    Dew point control and fuel conditioning exist to knock out heavy hydrocarbons from the gas stream. These operations are primarily field operations. Low Temperature Separators
    Low temperature separators (LTS) (also called low temperature extraction units, or LTX) are used both onshore and offshore. The process consists of cooling and partial condensation of the gas stream, followed by a low temperature separator.
    When inlet pressures are high enough to meet discharge-pressure requirements to make pressure drop acceptable, cooling is obtained by expansion through a J-T valve or turboexpander. Otherwise, external cooling is required. Water usually is present, and to prevent hydrate formation the separator downstream of the expander is warmed above the hydrate-formation temperature to prevent plugging. An alternative to heating is injection of either ethylene glycol or methanol, which is then recovered and dried for reuse.
    Figure 7.5 shows a LTS that uses ethylene glycol injection for hydrate prevention and uses J-T expansion for cooling. The feed initially goes through a water knockout vessel to remove free water. The water-saturated hydrocarbon gas and liquids then mix with ethylene glycol before being precooled. The mixture then passes through a J-T expansion valve and flashes into the low temperature separator to separate the gas, condensate, and glycol−water phases. The condensate goes to the condensate stabilizer for removal of remaining light ends.
    Overhead gas from the low temperature separator passes through the precooler before being combined with the stabilizer overhead and put into the pipeline. The low temperature separator is set to maintain the proper dew point of the blended outlet gas.
    The C3+ condensate from the stabilizer goes to product storage. The glycol−water mixture from the low temperature separator goes to the glycol regenerator for removal of the water and then reinjection into the feed.
    If inlet pressures are too low for expansion, the stream is cooled by propane refrigeration. The advantage of direct refrigeration is that the pressure drop is kept at a minimum. Hydrate formation must be considered with either feed dehydration upstream of the unit or inhibitor injection. Glycol injection is usually the more cost effective, but if used, it increases the required refrigeration duty. Twister
    Twister, is a new device (from 1997) used for dew point control and dehydration. Figure 7.6 shows a cutaway view of the device and denotes the salient parts of the unit. Gas enters and expands through a nozzle at sonic velocity, which drops both the temperature and pressure and causes droplet nucleation. The two-phase mixture then contacts a wing that creates a swirl and forces separation of the phases by centrifugal force. The gas and liquid are separated in the diffuser; the liquid is collected at the walls and dry gas exits in the center.
    Advantages of the system include:
    Simplicity. No moving parts and no utilities required.
    Small size and low weight. A 1-inch (24-mm) throat diameter, 6 feet (2 m) long tube can process 35 MMscfd (1 MMSm3/d) at 1,450 psia (100 bar).
    Driven by pressure ratio, not absolute pressure.
    Relatively low overall pressure drop. System recovers 65 to 80% of original pressure.
    High isentropic efficiency. Efficiency is around 90% compared with 75 to 85% for turboexpanders.
    Drawbacks of the system include:
    Requires a clean feed. Solids erode the tubing and wing, necessitating an inlet filter separator.
    Limited turndown capacity. Flow variability is limited to +/- 10% of designed flow. This limitation is mitigated by use of multiple tubes in parallel.
    Liquids exit with a slip gas. This mixture is typically 20 to 30% of the total flow volume. The mixture can go to a gas−liquid separator for recovery of the gas, which may require recompression.

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

    Fig. 7.6 Cutaway view of Twister device. (Courtesy of Twister BV.) Vortex Tube
    Vortex tubes use pressure drop to cool the gas phase but generate both a cold and warm gas stream. If streams are recombined, the overall effect is comparable to a J-T expansion. The principle of operation is the Ranque-Hilsch tube, developed in the 1940s and commonly marketed as a means to provide cold air from a compressed air stream.
    For dew point control, and dehydration, the device has the vortex tube and a liquid receiver connected to the tube. Gas enters the tube tangentially through several nozzles at one end of the tube, expands, and travels spirally at near sonic velocities to the other end. As it travels down the tube, warm and cool gas separate. The cool gas goes into the center of the tube. Warm gas vents in a radial direction at the end, but the cool gas is reflected back up the tube and exits just beyond the inlet nozzles.
    Condensation occurs in the cool gas, and the liquid is moved to the walls by centrifugal force, where it collects and drains into the receiver below. The overall cooling effect is comparable to that of a J-T expansion, with a low-temperature separator. However, the vortex tube combines the expansion and separation into a single step. The working pressure of the tube is 500 to 3,050 psig. The turndown ratio is 15% for a single tube but can be increased by use of multiple tubes in parallel; the optimum pressure drop is 25 to 35%. The vendor states that liquid condensation must be less than 10 wt%. The device performs well with up to 5% liquids in the inlet stream.
    The device has been used to dehydrate gas from underground storage. To prevent hydrate formation in the cold stream, TEG is added. Like Twister, the vortex tube has the advantage of simplicity and light weight. It could be useful where limited turndown is acceptable. It will be of most value when no compression is required. Membranes
    As discussed in Chapter 6, membranes are being used in several areas of gas processing, including dew pointing. Membranes are ideal for this application, provided preconditioning is adequate to protect the membrane, and little penalty exists for permeate compression. Figure 7.7 shows the flow configuration. Gas enters the membrane on the discharge side of the compressor, and the residual gas provides fuel to the compressor engine or turbine. The low pressure permeate is recycled to the suction for recompression to recover the permeate. Table 7.3 provides results for one field unit. Gas rates are low because only a slip stream needs to be processed for fuel.
    Like the previous two technologies, the process is simple and requires no moving parts. It too has the advantage of being relatively small and light weight. The technology is used on several offshore installations. Unlike the Twister and vortex tube, membranes have the advantage of a turndown ratio down to 50%, with no performance penalty. This property may not be an advantage for fuel gas conditioning, where flow rates should be stable.
    The table points out the selectivity of the membrane and may be poorer than that of the above two technologies. In fuel conditioning, the selectivity is not a major issue because of the relatively small fraction of gas that needs to be recompressed, and the enriched stream is recycled without requiring additional compression.
    Chapter 6 points out that membrane permeability is the product of the solubility and diffusion coefficient. For separation of light gases, the primary mode of selectivity is the diffusion coefficient. These membranes are silicone rubber compounds that preferentially absorb the heavy components.

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

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

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

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

    Table 7-3. Operating Conditions & Composition of NG Stream Using Membrane for Fuel Gas Conditioning Cooling by Expansion or External Refrigeration
    A general rule is to assume that ethane recovery increases with increased richness of the gas. This assumption is made because the ethane content in the vapor at the top of the column is set by column feed composition, along with temperature and pressure.
    At constant pressure and temperature, the ethane concentration in the liquid decreases with increasing C3+ fraction, which lowers the ethane concentration in the vapor and, thus, increases the percent ethane recovered. (However, this outcome will not always be the case in plants that use J-T or turboexpanders, because leaner gas puts less of a load on the refrigeration-expander system and may lower column temperatures and increase recovery).

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

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

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

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

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

    Fig. 7.9 Refrigerated lean oil absorption process.

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

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

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

  12. Re: Basics of Gas Field Processing Book "Full text"

    the book is not available in pdf format
    You can use the book here as a text
    you can purchase electronic or paper copy from Amazon
    Best regards


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