16-1

Gas processing covers a broad range of operations to pre-pare natural gas for market. Processes for removal of contami-nants such as H2S, CO2 and water are covered extensively in other sections of the Data Book. This chapter will cover the pro-cesses involved in recovering light hydrocarbon liquids either for sale when their value as liquids is higher than their value as gas components or they must be removed to avoid condensation. The equipment components included in the processes described are covered in other sections of the Data Book. This section will bring those components together in process configurations used for liquid production.

INTRODUCTIONThe recovery of light hydrocarbon liquids from natural gas

streams can range from simple dew point control to avoid liquid formation to deep ethane extraction. The desired degree of liq-uid recovery has a profound effect on process selection, com-plexity, and cost of the processing facility.

The term NGL (natural gas liquids ) is a general term which applies to liquids recovered from natural gas and as such refers to ethane and heavier products. The term LPG (liquefied petro-leum gas) describes hydrocarbon mixtures in which the main components are propane, iso and normal butane, propene and butenes. Typically in natural gas production olefins are not present in LPG.

Typically, modern gas processing facilities produce a single ethane plus product (normally called Y-grade) which is often sent offsite for fractionation and processing. Whether accom-plished on-site or at another facility, the mixed product will typically be fractionated to make products such as purity eth-ane, ethane-propane (EP), commercial propane, isobutane, nor-mal butane, mixed butanes, butane-gasoline (BG), and gasoline (or stabilized condensate). The degree of fractionation and the liquid products is market and geographically dependent.

Early efforts in the 20th century for liquid recovery involved compression and cooling of the gas stream and stabilization of a gasoline product. The lean oil absorption process was developed in the 1920s to increase recovery of gasoline and produce products with increasing quantities of butane. These gasoline products were, and still are, sold on a Reid vapor pressure (RVP) specifica-tion. Vapor pressures such as 69 or 83 kPa (abs) are common specifications for gasoline products. To further increase produc-tion of liquids, refrigerated lean oil absorption was developed in the 1950s. By cooling the oil and the gas with refrigeration, the absorber vapor outlet is leaner and propane product can be recov-ered. With the production of propane from lean oil plants, a mar-ket developed for LPG as a portable liquid fuel.

In lieu of using lean oil, refrigeration of the gas can be used for propane and heavier component recovery. The use of straight refrigeration typically results in a much more economical pro-cessing facility than using lean oil. The chilling of the gas can be accomplished with mechanical refrigeration, absorption refrig-eration, expansion through a J-T valve, or a combination. In order to achieve still lower processing temperatures, cascade

refrigeration, mixed refrigerants, and most significantly turbo-expander technologies have been developed and applied. With these technologies, recoveries of liquids can be significantly in-creased to achieve deep ethane recoveries. Early ethane recov-ery facilities targeted about 50% ethane recovery. As processes developed, ethane recovery efficiencies have increased to well over 90% in well integrated facilities.

In some instances heavy hydrocarbons are removed to con-trol the hydrocarbon dew point of the gas and prevent liquid from condensing in pipeline transmission and fuel systems. In this case the liquids are a byproduct of the processing and if no market exists for the liquids, they may be used as fuel. Alterna-tively, the liquids may be stabilized and marketed as conden-sate.

GAS COMPOSITIONThe gas composition has a major impact on the economics of

NGL recovery and the process selection. In general, gas with a greater quantity of liquefiable hydrocarbons produces a greater quantity of products and hence greater revenues for the gas processing facility. Richer gas also entails larger refrigeration duties, larger heat exchange surfaces and higher capital cost for a given recovery efficiency. Leaner gases generally require more severe processing conditions (lower temperatures) to achieve high recovery efficiencies and incur a higher cost per unit of liquid product.

Gases are typically characterized by the cubic meters of re-coverable hydrocarbons per thousand cubic meters of gas. This is commonly expressed as “liquid content.” Liquid content was traditionally meant to apply to propane and heavier compo-nents but is often used to include ethane. The liquid content of a gas can be calculated as shown in Example 16-1.

The other major consideration in the evaluation of NGL re-covery options is the specification of the residue sales gas. Sales specifications are usually concerned with a minimum Higher Heating Value (HHV) of the gas, but in some instances the maximum HHV can also be a consideration. The calculation of HHV is covered in Section 23 and in more detail in GPA Stan-dard 2172, “Calculation of Gross Heating Value, Relative Den-sity, and Compressibility Factor for Natural Gas Mixtures from Compositional Analysis.” In addition, for some gas sales the maximum and mininum Wobbe Number of the gas may be spec-ified. For more information on the calculation of Wobbe Num-ber, See Section 1 definitions.

Removal of liquids results in gas “shrinkage” and reduction of the HHV. This shrinkage represents a loss of revenue for the gas sales which must be considered in the economics of an NGL recovery plant. In general, sales gas specifications set the mini-mum HHV at 35.4-37.3 MJ/m3. Thus, if any components such as nitrogen or CO2 are present in the gas, sufficient ethane and heavier components must remain in the gas to meet the heating value specification. If little nitrogen or CO2 is present in the gas, the recovery level of the ethane and heavier components is then limited by markets, cost of recovery, and gas value. The

SECTION 16

Hydrocarbon Recovery

16-2

calculation of HHV and shrinkage cost is illustrated in Example 16-1.Example 16-1 — Find the liquid content of the gas mixture in Fig. 16-1. Find the HHV of the feed gas and the HHV of the residue gas with the following NGL recovery efficiencies: C2 – 90%, C3 – 98%, iC4/nC4 – 99%, C5+ – 100%. What is the shrink-age cost at $3.80/MkJ?

Solution Steps: Solution is shown in Fig. 16-1. From Fig. 23-2 obtain the m3/kmol for each of the compo-

nents. Multiply the mole fraction of each component (mole %/100) by the total moles of inlet gas and divide by 0.02369 km3/kmol to get the m3 liq/1000 m3 gas of each component. The total m3 liq/1000 m3 gas from Fig. 16-1 is 0.4169.

For the recoveries specified the net m3/day and residue com-position can be found as shown in Fig. 16-1. In order to compute the HHV of the two streams, the HHVs of each component are found in Fig. 23-2. Multiplying the individual HHVs by the mole % gives a total HHV of 41.628 for the feed gas and 36.260 for the residue gas.

The shrinkage volume can be found by the difference of the volume of the feed gas times the HHV and the volume of residue gas times its HHV. This volume is then multiplied by $3.80/MkJ to get the shrinkage value of the NGLs.

Shrinkage Value = [(9.34 • 41.628) – (8.38 • 36.260)] • $3.80/MkJ = $322 800/dayThe value of the NGLs in $/m3 versus the value of the com-

ponents in the residue gas in $/m3 or the “spread” between these values is the primary economic criteria for NGL recovery proj-ect evaluations.

LIQUID CONTENT CALCULATION

Component Feed Gas Mole % m3/kmol

Available Estimated Recovery % Net m3/day Residue Gas

Mole %m3/1000 m3 m3/dayN2 1.000 1.115CO2 3.000 3.346C1 85.000 94.808C2 5.800 0.08445 0.2077 1936 90 1742 0.647C3 3.000 0.08686 0.1102 1030 98 1009 0.067IC4 0.700 0.10322 0.0306 285 99 283 0.008NC4 0.800 0.09950 0.0337 315 99 311 0.009IC5 0.300 0.11552 0.0147 137 100 137 0.000NC5 0.200 0.11433 0.0097 90 100 90 0.000C6+ (Set as nC6 for example)

0.200 0.12980 0.0110 103 100 103 0.000

Total 100.000 0.4169 3896 3675 100.000MSm3/day 9.34 8.38SHRINKAGE CALCULATION

Component Feed Gas Mole%

Residue Gas Mole % HHV MJ/m3 Feed Gas

MJ/m3Residue Gas

MJ/m3

N2 1.000 1.115 0.0 0.00 0.00CO2 3.000 3.346 0.0 0.00 0.00C1 85.000 94.808 37.707 32.051 35.749C2 5.800 0.647 66.067 3.832 0.427C3 3.000 0.067 93.936 2.818 0.063IC4 0.700 0.008 121.404 0.850 0.010NC4 0.800 0.009 121.792 0.974 0.011IC5 0.300 0.000 149.363 0.448 0.000NC5 0.200 0.000 149.656 0.299 0.000C6+ 0.200 0.000 177.554 0.355 0.000Total 100.000 100.000 41.628 36.260MSm3/day 9.34 8.38

FIG. 16-1

Solution to Example 16-1

16-3

DEW POINT CONTROLRetrograde condensation has long been known to occur at

reservoir conditions. Recognition that it also occurs in typical processing conditions was an early result of computer calcula-tions using equations of state to predict vapor-liquid behavior. The phenomenon is illustrated in Fig. 16-2 showing dew point calculations for a gas stream leaving a separator at 38°C and 7000 kPa (abs). These dew point curves show that as the pres-sure is reduced, liquid is formed. The heavier the hydrocarbon, the more the dew point temperature increases as the pressure is lowered. The cricondentherm of the dew point curve is primarily determined by the presence of the heaviest component in the gas rather than just the total quantity of the heavy component in the feed gas so accurate determination of the heaviest com-ponents is important in establishing the phase envelope.

When gas is transported in pipelines, consideration must be given to the control of the formation of hydrocarbon liquids in the pipeline system. Condensation of liquid is a problem in me-tering, pressure drop and safe operation. Condensation of liquid can also be a major problem due to two-phase flow and liquid slugging.

To prevent the formation of liquids in the system, it is neces-sary to control the hydrocarbon dew point below the pipeline

operating conditions. Since the pipeline operating conditions are usually fixed by design and environmental considerations, single-phase flow can only be assured by removal of the heavier hydrocarbons from the gas.

Low Temperature SeparationSeveral methods can be used to reduce the hydrocarbon dew

point. If sufficient pressure is available, the removal can be ac-complished by expansion refrigeration in an LTS unit. The ex-pansion refrigeration system uses the Joule-Thomson effect to reduce the gas temperature upon adiabatic expansion. This temperature reduction results in not only hydrocarbon liquid condensation but also water condensation. The water is gener-ally removed as hydrates in this process, melted and removed. Thus, the process can actually accomplish dew point control of both water and hydrocarbon in a single unit.

Fig. 16-3 shows one example of an LTS system. The high pressure gas may first go through a heater but this is often not needed, depending on the gas conditions. The gas then enters the heat exchanger coil in the bottom of the separator where the gas is cooled by exchange with the condensed liquid and by melting the hydrates formed. Any water or condensate pro-duced at this point is removed in the high pressure separator. The gas from the separator is then heat exchanged with the

FIG. 16-2

Typical Low Pressure Retrograde Condensation Dewpoint Curves3

16-4

outlet product gas for further cooling. The temperature must be controlled at this point to prevent hydrate formation in the ex-changer. The gas from this point passes through the pressure reducing valve where the Joule-Thomson expansion occurs. The hydrocarbon liquid and hydrates produced from this expansion fall to the bottom of the low temperature separator. The hy-drates are melted and both the water and condensate are re-moved by level control. The gas leaving the separator has a hy-drocarbon dew point equal to the temperature and pressure of the separator.

The hydrocarbon and water dew points achievable with this process are limited by the pressure differential available as well as the composition of the feed gas. The LTS system can only be used where sufficient pressure is available to perform the de-sired processing and separation. It is an attractive process step if sufficient liquid removal can be achieved at the available op-erating conditions. A further modification to this process is to add glycol injection to the high pressure gas to allow the achieve-ment of lower water dew points when available pressure is lim-ited. Fig. 16-4 shows an LTS system with glycol injection. The use of the glycol eliminates the need to heat the LTS liquid phase and helps to ensure that no hydrate formation will block the process equipment upstream of the LTS.

The LTS may not work effectively if the removal of heavy hydrocarbon then changes the phase envelope so that partial condensation becomes infeasible. This is typical of lean gas containing small amounts of heavy hydrocarbons.

RefrigerationOften excess pressure is not available to operate an LTS sys-

tem. An alternative to the expansion refrigeration system is to utilize a mechanical refrigeration system to remove heavy hy-drocarbon components and reduce the gas dew point. The sche-matic for a refrigeration dew point control unit is shown in Fig. 16-5. This process flowsheet is essentially the same as that used for straight refrigeration NGL recovery. The gas pressure is generally maintained through the process allowing for equip-ment pressure drops. The gas is heat exchanged and then cooled by the refrigeration chiller to a specified temperature. Liquid is separated in the cold separator. The temperature of the separa-tor is set to provide the desired dew point margin for sales gas operations. This temperature specification must take into ac-count the gas which is recombined from the liquid stabilization step as well as potential variations in the feed gas pressure.

Provision must be made in this process for hydrate preven-tion. This can be accomplished by either dehydration upstream of the unit or by integrating the dehydration with the refrigera-tion unit. Use of glycol injection is usually the most cost effec-tive means of controlling water dew points. The only drawback is that the refrigeration must be in operation to accomplish the dehydration. If it is desired to operate the dehydration at times independent of the refrigeration, then separate units are used.

FIG. 16-3

Low-Temperature Separation Unit4

16-5

FIG. 16-4

Low-Temperature Separation System with Glycol Injection and Condensate Stabilization4

FIG. 16-5

Straight Refrigeration Process

16-6

StabilizationOne of the problems in using dew point control units of both

expansion LTS and mechanical refrigeration systems is the dis-position of the liquids removed. The liquids must be stabilized by flashing to lower pressure at high temperature or by using a stabilization column. When the condensate is flashed to a lower pressure, light hydrocarbons are liberated which can often be used in a fuel gas system.

The stabilization column can produce a higher quality and better controlled product. The condensate stabilizer is usually a top feed column which runs at a reduced pressure from the cold separator and has a reboiler to produce a specified vapor pressure product. The overhead vapor is either sent to fuel as shown in Fig. 16-4 or recompressed and combined with the sales gas as shown in Fig. 16-5. The column contains either trays or packing to provide necessary mass transfer for stabili-zation of the liquid feed. After stabilization, the product is cooled and sent to storage.

Compact Separation DesignsCompact process configurations have been commercialized

to take advantage of gas expansion for liquid separation. Each of these processes use static equipment to achieve the desired separation and are focused on replacing Joule-Thomson expan-sion valves and/or turbo-expanders.

One of these processes which is currently used in several operations is the Twister® technology. This process (Fig. 16-6) uses a supersonic nozzle in which the pressure is reduced via isentropic expansion such that liquid is formed. Saturated feed gas is passed across a vane ring which induces a swirling mo-tion where after the swirling gas is expanded in a Laval-type nozzle causing pressure and temperature to drop while the flow becomes supersonic. The centrifugal motion forces the liquid droplets to agglomerate at the wall from where it is drained

from the apparatus. The dry vapor stream is then decelerated in a diffuser duct to pipeline velocity, causing the pressure to increase to 70–80% of the feed pressure. Ample testing has shown that the Twister process has a high isentropic efficiency while separating almost all of the condensed liquid.

This technology is used for water and hydrocarbon dewpoint control in both onshore and offshore locations, especially where space and mass are at a premium. Since the expansion cooling in Twister is close to isentropic, more valuable NGL liquids are produced compared to a JT LTS process. A Twister based dew pointing process can often operate without glycol or methanol systems for hydrate suppression.

Another process uses a vortex tube device to affect the sepa-ration. The vortex tube is based on the Ranque-Hilsch tubes de-veloped in the 1940s. These tubes have been used as laboratory devices and small scale coolers. The working principle of these devices is the same. (Fig. 16-7) A gas is injected tangentially through a nozzle into the center of the tube where it expands to a low pressure. The gas flows cyclonic to the far end of the tube. During this flow, two temperature zones are formed, a warm zone near the wall and a cooler zone near the center. At the end of the tube the center gas is deflected and returns along the tube through an orifice near the inlet nozzle. The tube is therefore capable of producing two gas streams at different temperatures. The cold gas is at a temperature below that achievable with an isenthalpic expansion. If the two outlet streams were to be mixed, the combined temperature would be equal to the temperature achieved by the isenthalpic expansion. Thus the vortex tube per-forms the same function as the Joule-Thomson valve but pro-duces a lower outlet gas temperature for a portion of the stream. This apparatus could have application where gas pressure drop is available, dewpoint control is needed, and the warm and cool gases are recombined after liquid removal. This technology is very space efficient and can be attractive, especially for offshore processing applications.

FIG. 16-6

Concept of the Twister Process

16-7

Membrane ConditioningRubbery heavy hydrocarbon selective membranes are an-

other recent process development for use in fuel as condition-ing. These membranes are different from conventional cellu-losic membranes used for CO2 removal from natural gas as described in Section 21. These rubbery membranes preferen-tially permeate the heavy hydrocarbons through the membrane over methane, which is counterintuitive. The reason for this “reverse permeation” behavior is that the rubbery membranes separate gases based on solubility differences as compared to diffusion rate differences in the molecules.

This separation difference allows these membranes to be used for a variety of applications in the natural gas processing and treating areas. The main application of these rubbery mem-branes has been in remote area fuel gas conditioning to decrease the heating value of raw fuel gas. This process allows the raw gas to meet required engine specifications, thereby allowing op-erators to run compression driven by gas engines and move the gas from the remote location to the central processing facilities. One version of the membrane will also remove CO2/H2S as well as the heavy hydrocarbons. One company, MTR has dozens of units installed in various offshore, onshore remote locations, shale gas production areas and early production facilities all over the world4 (See Fig. 16-8).

Other commercial applications of the rubbery membranes are in nitrogen rejection (to reduce N2 content in natural gas), and for LPG/NGL recovery from gas streams. In the nitrogen rejection application, N2 is the least permeable component in the rubbery membrane, so the membrane can be used to reduce the N2 concentration in the pipeline gas, but in this case the methane rich gas will be in the low pressure permeate along with the heavier hydrocarbons. Therefore, the pipeline gas in this case would have a high heating value recovery, while re-jecting the nitrogen in the residue gas.

In LPG recovery, the fact that the heavy hydrocarbons per-meate the membrane, can be effectively used to increase the NGL content in the condensing loop, thereby allowing NGL re-covery at higher condensing temperature. NGL recovery mem-

brane processes are most suitable for flare gas applications in which NGL recovery is combined with power generation from the clean residue gas produced by the rubbery membrane.

FIG. 16-8

Commercial Conditions for Rubbery Membranes

Type of Rubbery Membrane

Pressure Max & (Range) kPa (ga)

Flow rateMax & (Range)

MSm3/dayStandard Hydrocarbon selective (SR Membranes)

8270 (345 – 8270) 3.4 (0.0028 – 3.4)

Polar Hydrocarbonselective membranes(Polaris Membranes)

6900 (345 – 6900) 0.28 (0.056 – 0.28)

Rotary Valve Fast Cycle PSA for Fuel Gas Conditioning

A fast cycle Pressure Swing Adsorption (PSA) developed by Xebec has been used successfully to condition gas.5 This process selectively adsorbs different gases by adsorbents when subject-ed to varying pressures. Inlet gas flows up through a vessel filled with a set of adsorbents at relatively high pressure. The heavier hydrocarbons (C2+), inerts (CO2 & N2) and water are preferentially retained, while methane is not. Before the bed becomes saturated, it is switched out of adsorbing mode, and the pressure is reduced. The lower pressure releases the impu-rities from the adsorbent as exhaust gases. The adsorbent bed is then re-pressurized and is ready to repeat the cycle.

To ensure a continuous supply of product gas, the process uses multiple adsorbent beds, which are cycled in a sequential manner through identical operating cycles. The system con-tains slowly rotating, multi-port, selector valves, which are cen-tral to the operation of the unit. These valves are used to create a very efficient, rapid PSA cycle, which results in smaller equip-ment packages with reduced capital and installation costs than a traditional PSA would require.

FIG. 16-7

Basic Design of a Vortex-Tube Device

16-8

STRAIGHT REFRIGERATION NGL RECOVERY

The straight refrigeration process is quite flexible in its ap-plication to NGL recovery. As outlined in the previous section, the process can simply be used for dew point control when mod-est liquid recovery is needed or desired. Alternatively, the pro-cess can be used for high propane recovery and, in the case of rich gases, for reasonable quantities of ethane recovery. The recovery level is a strong function of the feed gas pressure, gas composition and temperature level in the refrigeration chiller. Fig. 16-9 shows curves for estimating the recovery achievable as a function of temperature and gas richness for a given pro-cessing pressure. (Liquid content in this figure is propane plus.) Generally speaking, higher recovery efficiencies can be achieved with richer feed gas. The straight refrigeration process is typi-cally used with a glycol injection system. This configuration is limited in the temperature of operation by the viscosity of the glycol at the lower temperatures. Also, refrigeration is typically provided by propane refrigeration which is limited to –42°C re-frigerant at atmospheric pressure and thus a processing tem-perature of about –40°C. In order to go lower in processing tem-perature, upstream dehydration and alternative refrigeration systems must be considered.

Fig. 16-10 illustrates the ethane recovery efficiency which can be expected. As with propane recovery, for a given tempera-ture level, higher extraction efficiency can be achieved with richer gas. However, ethane recovery of over 30% can be achieved from a gas as lean as 0.4 m3/1000 m3 gas (C3+). Fig. 16-11 illustrates the effect of gas pressure on plant performance in propane plus recovery operation.

Refrigeration Process AlternativesThere are many variations in the straight refrigeration pro-

cess. Fig. 16-12 illustrates four of the most common variations. In the first scheme the gas is cooled against the residue gas and

FIG. 16-10

Recovery Efficiency, Ethane Plus5

FIG. 16-11

Effect of Gas Conditions on Propane Recovery5FIG. 16-9

Recovery Efficiency, Propane Plus5

16-9

FIG. 16-12

Refrigeration Process Alternatives6

16-10

the cold separator liquid before being chilled with refrigeration. This scheme uses a top-feed fractionator with the overhead be-ing recompressed and recycled to the inlet. The use of the liquid /feed gas exchanger helps reduce the chiller load. In this case, the residue gas from the cold separator has a dew point of the cold separator operating temperature.

The second scheme also uses a top-feed fractionator, but the cold separator liquid is fed directly to the fractionator. This fractionator operates with a lower overhead temperature which justifies exchange with the refrigeration system. The overhead after being warmed is recompressed and blended with the resi-due gas from the cold separator. In this configuration the frac-tionator overhead usually raises the residue gas dew point somewhat. The cold separator temperature must be set to en-sure that the desired dew point specification of the combined stream is achieved.

The third process uses a refluxed fractionator. This type of design usually has the highest liquid recovery efficiency, but has a higher cost due to the overhead reflux system. The fourth variation can be used where the cold separator liquid can be pumped and the stabilizer run at an elevated pressure. This can eliminate the need for a recompressor.

Any one or a combination of the following conditions:• Higher separator pressure• Richer gas• Recovery limited to propane-plus

will lead to higher recycle/recompression rates. This results in more refrigeration power, more recompressor power, more frac-

tionator heat, and larger equipment. These conditions favor the second and third schemes of Fig. 16-12.

Any one or a combination of the following conditions:• Lower separator pressure [around 4100 kPa (ga)]• Leaner gas (below 0.4 m3/1000 m3 gas C3+)• Recovery includes ethane

will lead to lower recycle/recompression rates. These conditions favor the first scheme in Fig. 16-12, or the fourth scheme if the separator pressure is not higher than 2750-3100 kPa (ga). Sep-arator pressure below 2750 kPa (ga), expecially with lean gas, will result in poor product recovery.

Regardless of the exact configuration employed, the capacity of the specific refrigeration system varies directly with refriger-ant condensing temperature and evaporating temperature. Condensing temperature is set by the condensing medium available at the plant site, and the process chiller temperature is set by the refrigerant evaporating temperature. Refrigerant power requirements vary with condensing and evaporating temperatures. Lower condenser temperature and higher evapo-rating temperature require lower power per unit of refrigera-tion required. For a given refrigeration load, power and con-denser duties can be found in Section 14 for a variety of refrigerants.

IPORSM ProcessThe Lummus Technology/Randall Gas Technologies licenses

the IPORSM process, a refrigeration based process, recovers 99.8+% C3 from gas streams with essentially no C2 recovery.

FIG. 16-13

IPORSM Process

16-11

This design incorporates a closed loop propane refrigeration cycle coupled with an open loop ethane mixed refrigeration cy-cle. This enables lower temperature operation and provides a reflux stream to the fractionation column, the combination of which produces high recovery for relatively low power consump-tion.

The process, shown in Fig. 16-13, can operate over a wide range of pressures, depending on the feed conditions and re-quired recoveries, and is well suited to handling rich feed gas. The optimum feed pressure is 1725 to 3790 kPa (ga). The pro-cess has an outlet pressure about 170 kPa lower than the inlet pressure, so recompression will be required to meet higher resi-due gas pressure battery limits.

The use of ethane-based refrigeration has been demonstrat-ed to provide efficient high liquids recovery, particularly when combined with simultaneous heat and mass transfer equip-ment.

LEAN OIL ABSORPTION NGL RECOVERY

Absorption is the physical process where a vapor molecule of a lighter hydrocarbon component will go into solution with a heavier hydrocarbon liquid (nonane, decane and heavier) and be separated from the gas stream. The process can be operated at ambient temperatures if only the heavier NGL products are desired. A refrigerated system enhances the recovery of lighter hydrocarbon products such as ethane and propane; lighter hy-drocarbon products such as ethane and propane; however, ethane recovery is limited to less than 50% to avoid having to remove excessive amounts of methane in the Rich Oil Demetha-nizer, which would lead to excessive solvent circulation. The absorbing fluid (lean oil) is usually a mixture of paraffinic com-pounds having a molecular mass between 100 and 200 assum-ing limited ethane recovery.

Lean oil absorption processes have the advantage that the absorber can operate at essentially feed gas pressure with min-imal loss of pressure in the gas stream which exits the process. Plants, whether ambient or refrigerated, are constructed of car-bon steel. This type of process was used from the early part of the 20th century and plants are still in use today. However,

most lean oil plants have been shut down or replaced with more modern straight refrigeration or turboexpander process plants. The lean oil process requires large processing equipment with excessive energy requirements. Lean oil absorption units are still used in many refinery operations.

Process ConsiderationsThe desired composition of the lean oil is determined by the

absorber pressure and temperature. The optimum molecular mass lean oil is the lowest mass oil which can be retained in the absorber with acceptable equilibrium losses to the residue gas. Lean oil absorption plants operating without refrigeration will require a higher molecular mass oil, usually in the 150-200 mo-lecular mass range. Refrigerated lean oil absorption systems can operate with an absorbing medium as low as 100 molecular mass with proper design.

Since the absorption is on a molar basis, it is desired to con-tact the gas stream with the maximum number of moles of lean oil to maximize the recovery of products from the gas. However, the circulation rate is in units of volume, cubic meters per hour. Therefore, a plant designed to circulate a heavier molecular mass oil can circulate more moles of oil with the same equip-ment if the molecular mass is lowered.

Many absorption oil recovery plants designed to originally operate at ambient temperatures have been modified to include a refrigeration system that allows both the lean oil and the gas to be chilled before entering the absorber. The reduced temper-ature increases the absorption and allows circulation of less oil of lower molecular mass because the vaporization rate into the residue gas is reduced. Oil is also lost with the NGL product. Oil losses with the product can be minimized by improving frac-tionation in the lean oil still. Many refrigerated lean oil absorp-tion plants can recover enough heavy ends from the gas stream to offset oil losses from the absorber, thereby making its own absorption oil.

If the gas stream contains compounds that cause the ab-sorption oil molecular mass to exceed design, a lean oil stripper can be used on a side stream of circulating lean oil to remove the heavy components. It is important to maintain the molecu-lar mass of the absorption oil at the design value because the circulating equipment, heat exchangers, and distillation pro-cess are designed to utilize a particular molecular mass fluid.

Refrigerated Lean OilFig. 16-14 shows a typical refrigerated lean oil absorption

process. The actual equipment configuration changes with dif-ferent gas feeds and product recoveries.

Raw gas enters the plant inlet separator upstream of the main process where inlet liquids are separated. The gas then enters a series of heat exchangers where cold process gas and the refrigerant reduce the feed gas temperature. This reduction in temperature results in condensation of the heavier hydrocar-bons in the inlet gas.

The gas is then fed to the bottom of the absorber where it flows upward countercurrent to the lean oil which is introduced at the top of the column. The lean oil has also been chilled to aid in NGL absorption. This column has trays or packing which increase the contact of the gas and lean oil. The lean oil physi-cally absorbs the heavier hydrocarbons from the gas. The light-er components stay in the gas and leave the top of the absorber. The oil and absorbed hydrocarbons leave the bottom of the ab-sorber as “rich oil.”

FIG. 16-14

Refrigerated Lean Oil Absorption8

16-12

The rich oil flows to the Rich Oil Demethanizer (ROD) where heat is applied to the rich oil stream to drive out the lighter hydrocarbons which were absorbed. Some of the cold lean oil is also fed to the top of the ROD to prevent loss of desirable NGLs from the rich oil.

The rich oil from the ROD is then fed to a fractionation tow-er or “still.” The still is operated at a low pressure and the NGLs are released from the rich oil by the combination of pressure reduction and heat addition in the still. The operation of the still is critical to the overall plant operation as this is not only the point where the desired product is produced, but the lean oil quality from the bottom of the column is important in the ab-sorption of NGLs in the absorber. The refrigeration required for the oil and gas chilling and the heat inputs to the ROD and still are the key parameters which must be controlled to operate a lean oil plant efficiently.

LOW TEMPERATURE NGL RECOVERY PROCESSES

Dew point control and mechanical refrigeration systems are intended for applications where enough liquefiable hydrocar-bons must be removed from the gas stream such that the gas can be transported without liquids formation and/or where some of the heavier components must be removed in order to meet a maximum heat content specification. In these cases, the goal is to meet the natural gas stream sales specification, not recovery of liquids. In cases where the liquefiable hydrocarbons present in a gas stream are more valuable as a liquid product than as the heating value they contribute to the gas, more effec-tive methods of liquids recovery are desired. The lean oil pro-cesses described in the last subsection were the original tech-nique used for this objective, but as already noted, this technology has been superseded by more cost effective cryogen-ic expansion process methods, which are covered in the follow-ing sub-sections.

In order to condense and separate more ethane and propane and thus achieve higher NGL recovery levels than are possible with dew point control or propane-based refrigeration alone, more gas cooling must be provided, which means cryogenic tem-peratures (below –45°C) are required. In order to achieve these temperatures, a combination of pressure expansion and chilling is used. The two types of cryogenic expansion processes in use today are:

1. Adiabatic J-T expansion across a control valve2. Isentropic expansion through a turboexpander3. Variants of each system above with supplemental me-

chanical refrigerationThe turboexpander processes are the most widely used;

however, J-T expansion type processes are sometimes used for lower recovery levels in unattended or remote locations, and usually for gas flow rates less than 0.566 MSm3/day. Turboex-pander type processes are used for higher ethane and propane recovery levels with larger gas inlet flow rates where the cost of the additional equipment is easily justified by the value of the additional liquids recovered.

One limitation on the recovery of ethane and heavier prod-ucts is the effect of liquids extraction on the heat content of the residue gas. The presence of inerts in the feed gas may limit the amount of ethane that can be removed and still meet the mini-mum heating value specification for the residue gas.

Another consideration is the liquids recovery level that is justifiable for a given gas composition and plant location. If eth-ane can be recovered at a good price differential over its value as a fuel in the residue gas, then the additional compression and complexity of a high ethane recovery facility may be justi-fied. If there is no local market for ethane, or no way to trans-port high vapor pressure product to market, then only propane and heavier components may have an economic incentive for recovery. The plant may be designed for initial operation in a propane recovery mode, with future conversion to ethane recov-ery if or when the local demand for ethane justifies the expense of additional compression. Many plants in North America are designed for dual mode operation in which the operator can se-lect high ethane recovery or ethane rejection, while still main-taining high propane recovery. These plants can be switched, during normal operation, as often as necessary between ethane recovery and ethane rejection, in response to changes in the processing margin. Historically, recovery of propane and heavi-er components is almost always justified.

Propane Plus NGL Recovery Compared to Ethane Plus NGL Recovery

Ethane is more volatile than the propane and heavier com-ponents in a natural gas stream so generally deeper cooling is required to condense the ethane fraction than is needed to con-dense only the propane and heavier components. The tempera-tures needed for ethane recovery are thus colder than those needed for propane recovery regardless of the process design. This means that more compression power is required for high ethane recovery than for propane recovery only because of the higher pressure differential required to produce lower tempera-tures across the cryogenic sections and/or refrigeration system.

The colder operating temperatures for ethane recovery mean that more of the plant will be constructed of stainless steel or aluminum than may be needed for a facility which only recovers propane. Another consequence of the colder tempera-tures is that the allowable concentration of CO2 in the feed to avoid CO2 freezing will be less for an ethane recovery design than for a propane recovery design.

As stated previously, many process designs can operate in either an ethane recovery mode or in an ethane rejection (pro-pane recovery) mode, as long as the equipment design tempera-tures are acceptable for both modes and the compression power is sufficient to support the desired ethane recovery level. An ethane recovery design will not be as efficient for propane re-covery compared to a propane recovery specific design. It is, however, possible to combine a very high efficiency ethane re-covery design with a very high efficiency propane recovery de-sign, if this need is identified during the design phase.9

There is always an incentive to minimize the power, emis-sions, and capital cost for the facility for a given recovery level of ethane and/or propane. A simple J-T expansion type process may have very good economic performance for a given loca-tion, flow rate and product price, but it will have low process recovery efficiency. This technology is therefore normally re-stricted to relatively small plants. Typical J-T cryogenic pro-cess configurations are described in the next sub-section. A good understanding of J-T operation will help in understand-ing the expander plant designs in the subsequent sub-sec-tions. Turboexpander plants are operated in a J-T mode before the turboexpander is started up and when the turboexpander is unavailable. Examples are provided later showing the dif-ference in product recovery for operation in J-T mode com-pared to turboexpander mode.

16-13

FIG. 16-16

J-T Process with Mechanical Refrigeration Recovery

INL E T GAS

NGL PR ODUC T

GAS

R E S IDUE C OMPR E S S OR

DE ME THANIZE R

J-T VAL VE

C OL D S E PAR ATOR

HE AT E XC HANGE

R E S IDUE

ME C HANIC AL

R E FR IGE R ATION

FIG. 16-15

J-T Process for Propane Recovery

INL E T G AS

L PG PR ODUC T

G AS

R E S IDUE C OMPR E S S OR

DE E THANIZE RJ-T VAL VE

C OL D S E P AR ATOR

HE AT E XC HANGE

R E S IDUE

16-14

INL E T G AS

FR OM DE HY

L PG PR ODUC T

G AS

R E S IDUE C OMPR E S S OR

DE E THANIZE RE XP ANDE R / C OMPR E S S OR

E XP ANDE R S E P AR ATOR

HE AT E XC HANGE

R E S IDUE

FIG. 16-17

Simple Turboexpander Process for Propane Recovery

DE E THANIZE R

INL E T G AS

FR OM DE HY

L PG PR ODUC T

G AS

R E S IDUE C OMPR E S S OR

E XP ANDE R / C OMPR E S S OR

E XP ANDE R S E P AR ATOR

HE AT E XC HANGE

S UB C OOL E R

R E S IDUE

FIG. 16-18

GSP Process for Propane Recovery

16-15

J-T EXPANSIONThe general concept for the Joule-Thomson (J-T) Expansion

design is to chill the gas by expanding the gas across a J-T valve. With appropriate heat exchange and large pressure dif-ferential across the J-T valve, cryogenic temperatures can be achieved resulting in acceptable extraction efficiencies. The main difference between the J-T design and turboexpander is that the gas expansion is adiabatic across the valve and is near-ly isentropic across the turboexpander. The outlet temperature is colder for a given pressure ratio for the turboexpander than for a J-T valve. The result is that the J-T design will have much lower ethane or propane recovery for a given amount of residue compression power than the turboexpander process, but the process is very simple, which helps make it suitable for some applications.

Specifically, the J-T process is used for situations where the higher efficiency due to the lower temperatures generated and the work recovered via the turboexpander driven compressor do not offset the increased cost & complexity involved, such as:

1. The feed gas rate is low and/or,2. There is a relatively low ethane and propane recovery re-

quirement3. The unit will be located in a remote or unattended loca-

tion4. Where a broad range of inlet gas flow rates and composi-

tions can be expected.

Typical Process Flow for J-T ProcessFig. 16-15 illustrates a typical process flow arrangement for

a J-T expansion process. In order to effectively use the J-T pro-cess, the gas must be at a high inlet pressure. Pressures over 6900 kPa (abs) are typical in these facilities. If the gas pressure is too low, inlet compression is necessary or sufficient expansion chilling will not be attained. The gas must first be dried to en-sure that no water enters the cold portion of the process. Typi-cally, molecular sieves or alumina are used for the drying. Methanol injection has been used in a few plants successfully but can be an operating problem because the dew point is too close to normal operating temperatures.

After drying, the gas is cooled by heat exchange with the cold residue gas and also by heat exchange with the demetha-nizer mid-tower liquids and in some cases the liquids from the cold separator. After chilling, the gas is expanded across the J-T valve and sent to the cold separator . The liquid from this sepa-rator is the feed to the demethanizer. This demethanizer col-umn is needed because the separator liquids contain too much methane and ethane to meet a liquid product specification for downstream fractionation. Usually this tower is a cold, top feed design. The cold liquid is demethanized to the proper specifica-tion in this tower. The cold overhead product from the demetha-nizer is used to cool the feed and is then recompressed as neces-sary for residue sales. When operated in an ethane recovery mode, enough heat is provided to the bottom reboiler to limit the methane content to an acceptable level. When operated in an ethane rejection mode, the bottom reboiler heat input is ad-justed to achieve an acceptable ethane content specification and the ethane is sent overhead and on to the residue gas sales point.

The liquids recovery for this process is determined by the pressure ratio across the J-T valve and the quantity of heat ex-change surface included in the plant heat exchangers. The pro-

cess can operate over a wide range of feed gas conditions and produce specification product. The process is thus very simple to operate and is often operated as an unattended or partially attended facility. The pressure ratio across the J-T valve is provided by the inlet gas pressure available upstream of the J-T valve and the column pressure downstream of the J-T valve. The power requirement for the residue gas compressor will in-crease as the column pressure is lowered. Higher relative resi-due compression power level results in higher pressure ratio across the valve and more cooling, so the recovery of the more volatile components increases. When the pressure ratio reaches 3:1, additional cooling can usually be supplied more efficiently using a refrigeration loop as described in the following sub-sec-tion. (Each megawatt of refrigeration will provide more incre-mental recovery than one megawatt increase in residue com-pressor power.)

Refrigerated J-T Expansion ProcessIn some cases the feed gas is not at high enough pressure or

the gas is rich in liquefiable hydrocarbons. Mechanical refrig-eration can be added to the J-T process to increase recovery of ethane and propane without adding a turboexpander. Fig. 16-16 shows a J-T process where a small packaged refrigeration system has been added to chill the feed gas. Another process variation is shown in this figure. The gas in this design is ex-panded downstream of the cold separator, taking advantage of the cooling provided by the refrigeration system and feeding the rich liquids separated at high pressure lower in the column. (The optimum location for the J-T valve is dependent on the inlet gas pressure and composition.) The advantage of exter-nal refrigeration is that lower feed pressure can be used or, alternatively, the demethanizer can be operated at a higher pressure thus reducing residue compression. For rich gas streams, using external refrigeration is more effective than in-creasing inlet or residue compression. The total compression power requirement (refrigeration plus residue plus inlet) may be minimized with the use of a refrigeration system although complexity is increased. The optimum combination of compres-sion power versus recovery level must be determined by process simulation of multiple equipment arrangements.

The J-T process, whether refrigerated or non-refrigerated, offers a simple, flexible process for relatively low ethane and propane recovery levels where a turboexpander is not justified. However, modern turboexpanders have proven to be so benefi-cial and reliable that almost all cryogenic plants are designed around turboexpander processes.

TURBOEXPANDER PROCESSESThe modern turboexpander based cryogenic gas plant was

commercialized in the mid-1960s. Since then almost all new high pressure NGL recovery plants built have employed a vari-ant of this basic process. Turboexpander process designs use the feed gas pressure to produce needed refrigeration by expan-sion across a turbine (turboexpander). The turboexpander re-covers useful work from this gas expansion. Typically the ex-pander is linked to a centrifugal compressor to recompress the residue gas from the process. Because the expansion is near is-entropic, the turboexpander lowers the gas temperature signifi-cantly more than expansion across a J-T valve. Details of the turboexpander equipment are in Section 13.

This extraction of energy from the inlet gas using a turboex-pander and its application to the booster compressor is why the inlet gas stream can be cooled more for a given residue compres-

16-16

sor power level than in a J-T process. More cooling means more condensation of the desired products so higher product recovery can be achieved for a given residue compressor size and power level.

PROPANE RECOVERY PROCESSESThe simplest (and least efficient) turboexpander process for

propane plus recovery uses a top feed, non-refluxed deethanizer column as shown in Fig. 16-17. Inlet gas is cooled then sepa-rated in to vapor and liquid phases at the cold separator. The vapor is routed to the expander inlet and the liquids are routed on level control through the inlet gas/liquids exchanger to a mid-column feed. The expander outlet is fed to the top of the deethanizer column where the vapor and liquids are separated. The vapor exits the column and the condensed liquids from the expander outlet are fractionated to separate the propane and heavier components desired in the bottoms liquids from the methane and ethane that were condensed as the vapor was cooled in the expander. Note that any components not con-densed in the expander outlet are lost to the residue gas stream. For a propane recovery mode plant, maximum practical recov-ery is limited to around 90% due to these losses at the top of the column.

The cold separator liquids also contain methane and ethane as well as propane and heavier components so the liquids are routed to a mid-column feed for fractionation as well. The col-umn bottoms are heated in a reboiler to a temperature suffi-cient to boil off the methane and ethane with minimum loss of propane to the overhead stream. The required bottoms temper-ature is always warmer than the plant inlet gas temperature so an external source of heat is needed for the reboiler. Hot oil, steam, or hot compressor discharge gas are commonly used for this service.

As higher ethane and propane recovery levels have been de-sired, more efficient NGL recovery designs have been devel-oped. The focus of these designs is to produce more reflux with leaner composition for the fractionation column so that the com-ponents not condensed at the expander outlet feed to the col-umn can still be recovered while reducing the compression power requirement.

GSP Design for Propane RecoveryThe configuration for a Gas Subcooled Process (GSP) design

as originally developed by Ortloff is shown in Fig. 16-18 below. The front end heat exchange arrangement is identical to the simple turboexpander process but the deethanizer column is ex-tended and stages are added above the expander feed. Reflux for this new column section is provided by taking a portion of the expander feed separator vapor and condensing and subcooling it at inlet pressure prior to let down on flow control to the column top feed. The reflux stream is very effective in condensing much of the propane that was not condensed in the expander outlet. The GSP design in propane recovery (ethane rejection) mode pro-vided a much needed improvement in efficiency over the simple turboexpander design,10 but still has lower efficiency than de-signs developed specifically for propane recovery.

Improved Open Art Propane Recovery Processes

The simple turboexpander process and GSP design described above are not as efficient for propane recovery as any of the proc-ess designs developed specifically for high propane recovery (ethane rejection). The simplest high propane recovery process is

an overhead recycle (OHR) design shown in Fig. 16-19. This process configuration uses an absorber column and a separate deethanizer fractionation column to achieve the desired recov-ery. The overhead from the deethanizer is condensed and used to absorb propane from the expander outlet stream. The expander outlet stream is fed to the bottom of the absorber column. As the reflux stream is warmed and partially vaporized by the expander outlet vapor, the vapor is cooled and additional propane is con-densed. The absorber overhead stream is colder than either of the feeds due to the absorption effect. The maximum propane recovery is limited by vaporization of propane from the reflux stream. This configuration provided much more efficient recov-ery of propane than the original simple turboexpander design.

Example Comparison for Propane Recovery Designs

The hypothetical example shown in Fig. 16-20 is presented to demonstrate the differences in the NGL recovery designs for propane plus recovery at constant residue compression power. The results for recovery of propane and heavier components us-ing J-T expansion, simple turboexpander, GSP, and overhead recycle processes are presented. Also tabulated are typical col-umn pressures and temperatures. The purpose here is to dem-onstrate the differences in operating conditions and power re-quirements for different processes.

ETHANE RECOVERY PROCESSESThe simple turboexpander process was originally used for

ethane recovery. When used for ethane recovery, the column pressure is lowered to provide a higher expansion ratio across the expander, thus generating more cooling and condensing more of the ethane in the expander outlet stream but with in-creased residue gas compression. The amount of heat needed in the bottom of the column is substantially less than for the eth-ane rejection mode because only the methane must be vapor-ized from the column bottoms liquids. The resulting column bottoms temperature is usually cool enough so that inlet gas can be used to supply some if not all of the required column duty. A column side heater may also be used to provide more inlet gas cooling while distributing the total heat input more uniformly to the column.

The simple turboexpander process for ethane recovery is shown in Fig. 16-21. Note the similarity to the J-T process with the J-T valve replaced by the expander. Dry feed gas is first cooled against the residue gas and used for side heating of the demethanizer. Additionally, with richer gas feeds, mechanical refrigeration is often needed to supplement the gas chilling. The chilled gas is sent to the cold separator where the con-densed liquid is separated, flashed and fed to the mid-point feed of the demethanizer. The vapor flows through the turbo ex-pander and feeds the top of the column. A J-T valve is always installed in parallel with the expander. This valve can be used for the plant inlet gas flow if the expander is out of service. When the expander is out of service, the plant operates like the J-T plant described previously, so the ethane recovery decreas-es substantially.

In this configuration the maximum ethane recovery is lim-ited to about 80%. The cold separator must be operated at a low temperature to maximize recovery. Often the high pressure and low temperature conditions are near the critical point of the gas making operation unstable when very small changes in separa-tor temperature result in large changes in the vapor volume to the expander. Another problem with this design is if there is more than a few tenths of a percent CO2 in the feed. The CO2

16-17

DE E THANIZE R

INL E T G AS

FR OM DE HY

L PG PR ODUC T

G AS

R E S IDUE C OMPR E S S OR

E XP ANDE R / C OMPR E S S OR

E XP ANDE R S E P AR ATOR

HE AT E XC HANGE

R E S IDUE

C ONDE NS E R

R E FL UX P UMPS

AB S OR B E R

FIG. 16-19

Overhead Recycle Process for Propane Recovery

A B C DProcess Design J-T SIMPLE GSP OHRCalculated Ethane Recovery % 0.20% 0.67% 0.75% 0.82%Calculated Propane Recovery % 22.8% 77.1% 86.4% 94.8%Calculated Butanes Recovery % 55.5% 97.5% 98.4% 99.8%

Total Liquids Recovered m3/day 1546 2146 2191 2266Expander Power kW 0 2386 3145 3667Column Overhead Temperature °C –28 –64 –69 –75Column Bottoms Temperature °C 160 116 109 107Column Reboiler Duty MW 3.0 4.5 5.1 5.6Residue Gas Flow Rate MSm3/day 6.88 6.71 6.68 6.68

Notes:

1. C2/C3 Ratio set at 2.0 mol% all designs

2. Inlet Gas Flow 7.1 MSm3/day, 6.6% Ethane, 2.8% Propane, 43°C, 6900 kPa (ga)

3. Residue Delivery 49°C, 6900 kPa (ga)

4. Residue Compressor Power set at 8799 kW

FIG. 16-20

Process Design Comparison for Propane Recovery

16-18

FIG. 16-21

Simple Turboexpander for Ethane Recovery

INL E T GAS

NGL PR ODUC T

GAS

R E S IDUE C OMPR E S S OR

DE ME THANIZE RE XPANDE R / C OMPR E S S OR

E XPANDE R S E PAR ATOR

HE AT E XC HANGE

R E S IDUE

FIG. 16-22

GSP Process for Ethane Recovery

DE ME THANIZE R

INL E T G AS

FR OM DE HY

NGL PR ODUC T

G AS

R E S IDUE C OMPR E S S OR

E XP ANDE R / C OMPR E S S OR

E XP ANDE R S E P AR ATOR

HE AT E XC HANGE

S UB C OOL E R

R E S IDUE

16-19

can solidify at operating temperatures found in the top stages of the demethanizer. Later in this Section under “Solid Forma-tion” there is more information on the CO2 freezing problem and the methods for prediction of solid CO2 formation.

GSP Design for Ethane RecoveryThe Gas Subcooled Process (GSP) was developed to over-

come the problems encountered with the conventional expander process for ethane recovery. This process, shown in Fig. 16-22, deviates from the original expander process in several ways. A portion of the gas from the cold separator is sent to a heat ex-changer where it is totally condensed and subcooled at inlet gas pressure using the cold column overhead stream. This stream is then flashed across a flow control valve to the top feed of the demethanizer providing reflux to the demethanizer. Because the stream is subcooled, only a small fraction of the stream flashes at the reflux feed point.

The expander feed is sent to the tower several stages below the top of the column. Vapor rising from the expander feed will contain a significant amount of ethane, which is condensed by the colder reflux stream.

In this process the column overhead is warmed up and the column pressure is increased significantly without sacrificing liquid recovery, due to the use of the subcooled reflux and the cold separator now operates at a much warmer temperature, well away from the system critical. The flow rate through the expander is less than in a non-refluxed design, but the vapor is much warmer, so the expander power is actually higher for the GSP design than for a non-refluxed design. The residue com-

pression power is always much less at a given ethane recovery level than for the non-refluxed design so the non-refluxed de-sign is no longer used.

Example Comparison of Ethane Recovery Designs

The hypothetical example presented earlier for propane re-covery is presented in Fig. 16-23, but for ethane recovery to demonstrate the differences in the ethane recovery for several ethane recovery designs. The design basis assumptions from the propane plus recovery comparison are used here with the additional constraint of 2.0% max methane/ethane ratio for the bottoms product. For ethane recovery, only the J-T, simple tur-boexpander, and GSP options are tabulated since the overhead recycle process is a propane recovery design.

LICENSED NGL RECOVERY PROCESSES

Most ethane recovery plants today are screened initially us-ing a GSP design to determine rough power requirements and to see if refrigeration is needed. A GSP design may provide good results when the CO2 in the feed is low, the gas is not too rich, and only the ethane recovery mode of operation is desired. How-ever, there are many design basis requirements for which a li-censed design is advantageous. Many locations have require-ments that cannot be met with the process designs described above. Some of these needs can be met using licensed designs available from several sources.

FIG. 16-23

Process Design Comparison for Ethane Recovery

A B CProcess Design J-T SIMPLE GSPCalculated Ethane Recovery % 8.2% 59.3% 84.0%Calculated Propane Recovery % 27.1% 93.8% 98.7%Calculated Butanes Recovery % 56.5% 99.2% 99.8%

Total Liquids Recovered m3/day 1136 2753 3203Expander Power kW 0 3096 2820Column Overhead Temperature °C –29 –83 –94Column Bottoms Temperature °C 113 34 29Column Reboiler Duty MW 2.1 2.1 1.8Residue Gas Flow Rate MSm3/day 6.82 6.37 6.26

Notes:

1. C2/C3 Ratio set at 2.0 mol% all designs

2. Inlet Gas Flow 7.1 MSm3/day, 6.6% Ethane, 2.8% Propane, 43°C, 6900 kPa (ga)

3. Residue Delivery 49°C, 6900 kPa (ga)

4. Residue Compressor Power set at 9620 kW

16-20

FIG. 16-24

IOR® Propane Recovery Process

DE E THANIZE R

INL E T G AS

FR OM DE HY

L PG PR ODUC T

G AS

R E S IDUE C OMPR E S S OR

E XP ANDE R / C OMPR E S S OR

E XP ANDE R S E P AR ATOR

HE AT E XC HANGE

R E S IDUE

C ONDE NS E R

R E FL UX P UMPS

AB S OR B E R

A

DEETHANIZER

INLET GAS

FROM DEHY

LPG PRODUCT

GAS

RESIDUE COMPRESSOR

EXPANDER / COMPRESSOR

EXPANDER SEPARATOR

HEAT EXCHANGE

RESIDUECONDENSER

REFLUX PUMPS

SEPARATOR

FIG. 16-25

SCORE® Propane Recovery Process

16-21

IOR® and SCORE® Propane Recovery Processes

The basic overhead recycle process has been modified by Ortloff to make better use of the refrigeration available in the feed streams. This improved version known as IOR® process as shown in Fig. 16-24 makes a few improvements to the basic overhead recycle process. In this process the reflux for the deethanizer is produced in an absorber overhead system, which produces reflux for both towers. Adding this reflux to the deeth-anizer column helps minimize the amount of propane in the reflux thus providing a leaner reflux, so the column pressure can be higher for a given propane recovery. The absorber bot-toms are heated against the feed before being sent to the deeth-anizer thus cooling the inlet gas.

This process is typically designed for 99% propane recovery, while rejecting 98% of the ethane. The bottoms temperature is set to limit ethane in the bottoms to a level which will allow the propane product at the downstream fractionator to meet a va-por pressure specification. The improved heat integration and reflux utilization result in a 5-15% reduction in the residue compressor power requirement over the basic overhead recycle design at the same recovery level.

Ortloff later developed a simplified version of the IOR® proc-ess by combining the absorber and deethanizer into a single column with a vapor side draw below the expander feed to pro-duce reflux. This process shown in Fig. 16-25 is known as the Single Column Overhead Recycle (SCORE®) process.

IPSI Enhanced NGL Recovery ProcessAnother improvement of the turboexpander-based NGL

GSP design is the IPSI Enhanced NGL Recovery® Process11 shown in Fig. 16-26. This process utilizes a slip stream from or near the bottom of the distillation column (demethanizer) as a mixed refrigerant. The mixed refrigerant is totally or partially vaporized, providing refrigeration for inlet gas cooling other-wise normally accomplished using an external refrigeration system. The vapor generated from this “self-refrigeration” cycle is specifically tailored to enhance separation efficiency, then is recompressed and recycled back to the bottom of the tower where it serves as a stripping gas. The innovation not only re-duces or eliminates the need for inlet gas cooling via external refrigeration, but also provides the following enhancements to the demethanizer operation:

• Lowers the tower temperature profile, permitting better energy integration for inlet gas cooling via reboilers, re-ducing heating & refrigeration requirements.

• Reduces and/or eliminates the need for external reboiler heat, thereby saving fuel plus refrigeration.

• Enhances the relative volatility of the key components in the tower when operated at a typical pressure, improving separation efficiency and NGL recovery; or allowing in-creased tower pressure with lower recovery efficiency, reducing the residue gas compression requirements.

LPG-MAXSM ProcessThe Lummus Technology/Randall Gas Technologies LPG-

MAXSM process recovers 99+% propane from gas streams with essentially no ethane recovery. A simplified version of the scheme is available that has lower recoveries (93 to 95%) with lower capital cost.

This process, shown in Fig. 16-27, uses an absorber at high pressure to minimize recompression. The pressure of the resi-due gas leaving the booster compressor (driven by the expander) will depend on feed pressure, and recovery required. Process conditions are tailored to maximize pressure at booster outlet.

RSV® High Ethane Recovery ProcessThe limitation to ethane recovery for the GSP design is the

ethane content in the reflux liquids. The GSP reflux feed has the same composition as the cold separator vapor so it contains some ethane. A portion of this ethane will flash at the top feed, thus generally limiting the recovery to around 93%, regardless of the residue compressor power used. A leaner reflux stream is needed to achieve recoveries higher than 93%. By adding stages above the GSP reflux feed, and routing a condensed and sub-cooled stream, created by recycling a small portion of the resi-due gas back through the gas/gas exchanger and subcooler at residue gas delivery pressure, to provide a lean reflux stream, up to 99% ethane recovery can be achieved.

This Recycle Split Vapor (Ortloff RSV®) process design shown in Fig. 16-27 is an efficient process in very high ethane recovery designs, because the recycle rate is minimized by maintaining the GSP reflux stream at the second feed for bulk ethane recovery, and then only using enough recycle reflux at the top feed to capture the ethane from the equilibrium losses at the GSP reflux feed point.

NGL-MAXSM High Ethane Recovery ProcessThe Lummus Technology/Randall Gas Technologies NGL-

MAXSM process recovers 99+% ethane from gas streams with essentially 100% propane-plus recovery. This process, shown in Fig. 16-29, uses semi-lean and lean reflux to achieve very high ethane recovery or to increase gas throughput at lower recover-ies.

The process can be designed to run in propane recovery mode (dual operation, ethane recovery, and propane recovery). Switching from ethane recovery mode to propane recovery mode will ensure that there is almost complete ethane rejection with 99% C3+ recovery.

Other Licensed NGL Recovery ProcessesOther examples of licensed designs include the following:1. CO2 content in the feed too high for the liquid product

CO2/C2 ratio without having to reboil the column to a CO2 specification and lose ethane recovery. (Ortloff CDC)

2. CO2 content in the feed high enough to freeze using GSP design, but inlet treating is not justified. (Ortloff RSV)

3. Retrofit of an existing plant is limited by the existing in-let gas system throughput. (Technip DEER)

4. Dual mode and intermediate ethane recovery levels are desired without loss of propane recovery. (Ortloff SRP)

5. Dual column designs for high propane recovery with flex-ible ethane recovery (Lummus Randall Super Hy-Pro STC & TTC)

6. Integration of NGL Recovery and LNG production are desired. (Ortloff, IPSI)

16-22

7. Integration of NGL Recovery and Nitrogen Rejection are desired. (Costain/Chart) This type of design is discussed in more detail in the Nitrogen Rejection subsection.

AVOIDING COMMON OPERATING PROBLEMS

Most operating problems with cryogenic NGL recovery plants can be avoided with attention to detail in the design and attention to cleanliness during construction and maintenance. The following guidelines apply to both open art and licensed designs.

1. Mole sieve dust from the dehydrators will plug the brazed aluminum heat exchangers (BAHE) used in most cryo-genic plants. Very high quality adequately sized filters must be installed upstream of the heat exchangers. They must be installed in a two by 100% configuration without a bypass. Inlet gas must never be allowed to bypass the filters, and differential pressure must be monitored. The importance of good filtration and good operating practice cannot be overemphasized. Sudden plugging of the filters with mole sieve dust indicates disintegration of the mole sieve itself or failure of the bed supports and must be ad-dressed immediately to prevent permanent plugging of the exchangers.

2. The mole sieve dehydration system cooling gas tempera-ture should be cool enough to get the cooled dehy bed down to as close to the inlet gas temperature as possible. If the regenerated bed is too warm when placed into ser-vice, an inlet gas temperature spike will result as the bed is further cooled by inlet gas. The inlet gas will warm up and create a disturbance throughout the cryo plant. This disturbance will occur every time the dehys switch.

3. During commissioning, construction debris will typically migrate from the column packing or trays to the liquid draws from the column. The debris will then flow to the side heaters/reboilers and accumulate at the strainer. Two parallel 100% strainers with block valves on the liq-uid draw lines will allow cleaning of the strainers with-out a shutdown.

4. The initial design must include low point drains to allow blow down of all free water prior to cool down. Dryout lines from the dehys to the bottom of the column and to the reflux system are typically installed to allow a flow of warm dry gas through the entire plant to remove mois-ture. A plant recycle line from the residue compressor discharge back to the dehy system will allow dryout of the plant with minimum flaring. Dryout must be ad-dressed during the design phase of the project.

5. Changes in reflux flow rates or bottoms temperature set-points may take several hours to settle out in process de-signs where there are recycle type reflux systems. Opera-tors must allow sufficient time for the plant to stabilize when making even small changes as feed composition or flow rate changes.

6. Any decrease in pressure ratio across expander the ex-pander will reduce recovery. Any increase in pressure drop across strainers or exchangers will result in lower expander inlet pressure or higher column pressure and normal recovery will not be achieved until the normal pressures at the expander are reestablished. Gradual increases in pressure drop on the inlet gas side of the heat exchangers may indicate mole sieve dust getting through damaged filter elements. Gradual increases in pressure drop across the cold exchangers may indicate deteriorating dehydration system performance and a

FIG. 16-26

IPSI Enhanced NGL Recovery Process

16-23

DE ME THANIZE R

INL E T GAS

NGL PR ODUC T

GAS

R E S IDUE C OMPR E S S OR

E XPANDE R / C OMPR E S S OR

C OL D S E PAR ATOR

HE AT E XC HANGE

R E S IDUE

S UB C OOL E R / GAS-G AS

E XC HANGE R

FIG. 16-28

RSV® Ethane Recovery Process

FIG. 16-27

LPG-MAXSM Propane Recovery Process

16-24

FIG. 29

NGL-MAXSM Ethane Recovery Process

FIG. 16-30

Four-column Fractionation System

16-25

gradual accumulation of ice. Methanol injection is typi-cally used to determine if the pressure drop is due to ice or particulates. Note that if operating below –98°C, Methanol will freeze.

7. One treating requirement that becomes important to con-sider in design when using these processes is to confirm if mercury can be present in the feed gas, and if so to add treating facilities to ensure any mercury present is re-moved to very low levels. A typical specification will be <0.1 microgram per cubic meter (<0.01 ppbv)) of inlet gas. See Section 21 — Mercury Removal for more infor-mation).

8. When molecular sieves are regenerated species co-ad-sorbed on the sieve are also removed. This means that when the newly regenerated bed is brought back on line there will be a short period where the species co-adsorbed such as any acid gases still present and more important-ly, light hydrocarbons like ethane, propane and butane, are re-adsorbed, and these species are removed from the gas stream. If this will have a negative impact on the downstream NGL recovery process, then a gas contain-ing these species needs to be used as part of the regenera-tion process to pre-saturate the bed prior to it being placed back in drying service.

FRACTIONATION CONSIDERATIONSThe cryogenic plant fractionation column produces the spec-

ification C2+ or C3+ stream as a bottom product. This mixed product then needs to be separated into usable products in a series of one or more fractionation columns. This fractionation may take place at the cryogenic plant facility or at some frac-tionation facility down the pipeline. However, many gas plants have onsite fractionation of the recovered NGL stream so that ethane, propane, and butane can be sold separately from that site.10

If the NGL stream is an ethane plus stream, the first step is to separate the ethane from the propane and heavier compo-nents in a deethanizer. The propane is then separated from the butane and heavier components in a depropanizer. If further processing is desired the butane may be separated in a debuta-nizer and the butanes further separated in a butane splitter column. The butane splitter is only used when a differential value can be realized for the isobutane versus the mixed butane stream. A schematic of a four column fractionator is shown in Fig. 16-30. Section 19 in the Data Book covers the specifics of fractionation systems for NGL streams. Some markets will re-quire a slightly different fractionation system arrangement in which the deethanizer bottoms liquid is routed to an LPG col-umn where the C3 and C4 LPG components are separated from

FIG. 16-31

Nine-stage Cascade Liquefaction Process12

16-26

the C5+ components. The C3+C4 stream may then be further fractionated in to separate C3 and C4 components if the local LPG spec cannot be met without reducing the amount of one component in the LPG product and selling that component sep-arately.

LIQUEFIED NATURAL GAS PRODUCTION

The principal reason for liquefying natural gas is the 600-fold reduction in the volume which occurs with the vapor-to-liquid phase change. This volume reduction is important in the transportation and storage of the gas. In the liquid state, the gas can be transported in discrete quantities, can be economi-cally stored in tanks for use as required, and can be transported long distances not feasible with gas pipelines.

Because methane is the primary component of natural gas, the production of Liquefied Natural Gas (LNG) involves the chilling of the entire natural gas feed stream to cryogenic tem-peratures sufficient to totally condense the gas stream. Com-mon to all LNG liquefaction processes is the need to pretreat the gas to remove components, such as CO2 and water, which will solidify in the liquefaction step. The liquefaction unit also has to remove hydrocarbon components, such as benzene and cyclohexane, which can solidify. Two types of LNG facilities have been developed: 1) large base load units for continuous LNG production to export markets, and 2) small peak shaving plants for gas distribution systems. The large scale based load units are typically designed with emphasis on process efficien-cy. In addition to the process units involved in the liquefaction step, base load LNG plants tend to be large complex facilities which involve product storage, loading and complete stand-alone utility systems.

Peak shaving facilities differ from base load units in several aspects. Peak shaving plants are much smaller, operate only a portion of the year, and are often located near the point of use for the gas. The design emphasis is thus on capital cost minimi-zation rather than thermodynamic efficiency. In order to pro-duce the low temperature necessary for liquefaction, mechani-cal refrigeration systems are utilized. Four types of liquefaction processes can be used to accomplish this refrigeration:

1. Cascade Refrigeration Processes2. Mixed Refrigerant Processes3. Precooled Mixed Refrigerant Processes4. Turboexpander Based Process (Peak Shaving)Each of these processes has been used for liquefaction facili-

ties with the PreCooled process being the predominant technol-ogy in base load units.

As of 2010 in the United States these processes are all being used in peak shaving units with the following distribution.

Mixed Refrigerant: 41

Expander Based: 14

Cascade or Nitrogen: 5

Total: 60

Cascade RefrigerationThe first LNG liquefaction units utilized the cascade refrig-

eration process. In the classical cascade cycle three refrigera-tion systems are employed: propane, ethylene and methane. Two or three levels of evaporating pressures are used for each of the refrigerants with multistage compressors. Thus the re-frigerants are supplied at eight or nine discrete temperature levels.Using these refrigeration levels, heat is removed from the gas at successively lower temperatures. The low level heat removed by the methane cycle is transferred to the ethylene cycle, and the heat removed in the ethylene cycle is transferred to the propane cycle. Final rejection of the heat from the pro-pane system is accomplished with either water or air cooling.

Early facilities used a closed methane refrigeration loop. More modern designs use an open methane loop such as shown in Fig. 16-31 where the methane used for refrigerant is com-bined with the feed gas and forms part of the LNG product. The efficiency and cost of the process is dependent on the number of refrigeration levels provided in each refrigeration system.

The refrigeration heat exchange units traditionally were based on shell and tube exchangers or aluminum plate fin ex-changers. Newer designs incorporate plate fin exchangers in a vessel known as “core-in-kettle” designs. A critical design ele-ment in these systems is the temperature approach which can be reached in the heat exchangers.

Mixed Refrigerant ProcessesAfter initial developments of cascade LNG plants, the mixed

refrigerant cycle was developed to simplify the refrigeration system. This system uses a single mixed refrigerant composed of nitrogen, methane, ethane, propane, butane and pentane. The refrigerant is designed so that the refrigerant boiling curve nearly matches the cooling curve of the gas being liquefied. The closeness of the match of these two curves is a direct measure of the efficiency of the process.

The process (Fig. 16-32) has two major components: the re-frigeration system and the main exchanger cold box. The cold box is a series of aluminum plate fin exchangers which provide very close temperature approaches between the respective pro-cess streams. The low pressure refrigerant is compressed and condensed against air or water in a closed system. The refriger-ant is not totally condensed before being sent to the cold box. The high pressure vapor and liquid refrigerant streams are combined and condensed in the main exchanger. The condensed stream is flashed across a J-T valve and this low pressure re-frigerant provides the refrigeration for both the feed gas and the high pressure refrigerant.

Removal of pentane and heavier hydrocarbons from the feed gas is accomplished by bringing the partially condensed gas out of the cold box and separating the liquid at an intermediate temperature. The liquid removed is then further processed to produce a specification C5+ product. Light products from this separation are returned to the liquefaction system.

Precooled Mixed Refrigerant ProcessThe propane precooled mixed refrigerant process (Fig. 16-33)

was developed from a combination of the cascade and mixed re-frigerant processes. In this process, the initial cooling of the feed gas is accomplished by using a multistage propane refrigeration system. The gas is cooled with this system to around –40°C at which point the gas is processed in a scrub column to remove the

16-27

FIG. 16-32

Mixed Refrigerant Liquefaction Process13

FIG. 16-33

Propane Precooled Mixed Refrigerant Process14

16-28

heavy hydrocarbons. The gas is then condensed in a two step mixed refrigerant process. The chilling of the gas is accomplished in a single, large, spiral-wound heat exchanger. This exchanger allows extremely close temperature approaches between the re-frigerant and the gas to be achieved.

The mixed refrigerant in this process is a lighter mixture composed of nitrogen, methane, ethane and propane with a mo-lecular mass around 25. The mixed refrigerant after recompres-sion is partially cooled with air or water and then further cooled in the propane refrigeration system. The partially condensed refrigerant from the propane chilling is separated and the high pressure vapor and liquid streams sent separately to the main exchanger. The liquid is flashed and provides the initial chill-ing of the gas. The high pressure vapor is condensed in the main exchanger and provides the low level, final liquefaction of the gas. As in the other processes, the LNG leaves the exchang-er subcooled and is flashed for fuel recovery and pumped to storage.

Dual Mixed Refrigerant Cycle ProcessesBuilding on the success of the propane pre-cooled mixed re-

frigerant cycle Shell and Linde have both successfully commer-cialized a dual or double mixed refrigerant cycle process. In this process the multi-stage propane refrigeration system used to cool the inlet gas and partially condense the mixed refrigerant is replaced with a second “warm” mixed refrigerant tailored to provide this cooling.

Using this warm mixed refrigerant tailored to match the cooling curve down to the –46 to –51°C level and then a cold mixed refrigerant to meet the refrigeration requirements below that level allows this process to efficiently produce the desired LNG stream.

Fig. 16-34 shows a simple schematic of one version of this design, which has been used commercially in several locations for base load liquefaction.

Precooled with Nitrogen Cycle Liquefaction Process

APCI has developed the AP-X® liquefaction process, which uses a third refrigerant cycle (nitrogen) to unload a propane pre-cooled mixed refrigerant (C3MR) process. All of the 7.8 mil-lion MTA LNG trains constructed in the last six years in Qatar use this new hybrid process.

The front of the hybrid process as shown in Fig. 16-35 is es-sentially identical to that of a conventional C3MR process. As in that process, propane chills the inlet gas to around –40°C, and the mixed refrigerant cools the gas down to only –115°C in a spiral-wound heat exchanger (SWHE), instead of –151°C typi-cal of the C3MR process.

The nitrogen expansion cycle then serves as the subcooling refrigerant, taking the liquefied gas temperature down the rest of the way, cooling it to –151°C. As with the C3MR process, the rest of the temperature decrease occurs in the final pressure reduction step.

The nitrogen cycle enables an increase in the train’s through-put at a given SWHE diameter. The SWHE diameter is what normally limits the capacity of a liquefaction train using SWHE’s. By adding the nitrogen cycle, several refrigerant streams are removed from the SWHE relative to the C3MR process, making room for more tubes dedicated to feedstock flow, and shortening the SWHE. The trade-off is the need for a second SWHE in subcooling (nitrogen) service.

Turboexpander Based Process Another route for commercialization not used in current on-

shore base load facilities is to use a liquefaction scheme based on the use of isentropic expansion via a set of turboexpanders to supply the refrigeration. A version of this is used in some peak shaving applications where only a portion of the inlet gas has to be liquefied.

This type of technology, while not as efficient as currently used cascade and mixed refrigerant based processes has the ad-vantage of avoiding or minimizing the refrigerant inventory re-quired for the liquefaction step, and therefore, might be attrac-tive for use offshore where large volumes of refrigerant would be hazardous and lead to much greater mass and deck area15.

SOLIDS FORMATIONIn addition to the obvious need for water removal from the

gas stream to protect from blockage in the cryogenic sections of a plant, consideration must be given to the possible formation of other solids or semi-solids in the gas stream. Amines, glycols, and compressor lube oils in the gas stream can form blockages in the system. Generally these contaminants will form a block-age upstream of an expander, in the lower temperature ex-change circuit, or on the screen ahead of an expander.

Carbon dioxide can form as a solid in lower temperature sys-tems. Fig. 16-36 will provide a quick estimate for the possibility of formation of solid CO2. If operating conditions are in the methane liquid region as shown by the insert graph, the dashed solid-liquid phase equilibrium line is used. For other conditions the solid isobars define the approximate CO2 vapor concentra-tion limits.

For example, consider a pressure of 2400 kPa (abs). At –110°C, the insert graph (Fig. 16-36) shows the operating conditions to be in the liquid phase region. The dashed solid-liquid phase equilibrium line indicates that 2.1 mol percent CO2 in the liquid phase would be likely to form solids. However, at the same pres-sure and –100°C, conditions are in the vapor phase, and 1.28 mole percent CO2 in the vapor could lead to solids forma-tion. This chart represents an approximation of CO2 solid for-mation. Detailed calculations should be carried out if Fig. 16-33 indicates operation in a marginal range. This means that the most probable condition for solid CO2 formation may be several trays below the top of the tower rather than at expander outlet conditions. Again, while Fig. 16-36 indicates marginal safety from solids formation, detailed calculations must be carried out.

In addition to CO2 and water which can solidify and cause blockage and damage in cryogenic equipment, hydrocarbons can also solidify at temperatures found in LNG plants. Figure 16-34 shows some freezing point temperatures for pure com-pounds which can be troublesome in LNG facilities.

While all of the compounds listed can solidify at LNG tem-peratures, the solubility of these compounds in LNG are such that only at certain concentrations will there be solid formation. Cyclohexane and benzene are the compounds with the highest freezing points in this list. Cyclohexane is normally not present in significant quantity in produced gas, however; benzene is present in most gases and can be found in level in the 1000 ppm range, well above the solubility limit at LNG temperatures. Toluene can also be a problem, so typically aromatic compounds are the highest concern due to the combination of concentration present in the gas and freezing point.

16-29

FIG. 16-35

AP-X® LNG Process

FIG. 16-34

Typical Dual Mixed Refrigerant Process

16-30

FIG.16-36

Approximate Solid CO2 Formation Conditions

16-31

FIG. 16-37

Solubility of Benzene in Methane

16-32

FIG. 16-38

Solubility of Benzene in Ethane

16-33

The solubility of these compounds in LNG streams is compo-sition dependent. Fig. 16-37 shows the solubility of benzene in methane, Fig.16-38 shows the solubility of benzene in ethane. Comparison of these two figures shows that at say –162°C, the solubility in methane is about 2 ppm while the solubility in eth-ane is 75 ppm. Heavier hydrocarbons such as propane and bu-tane have even higher solubility numbers. Thus the composi-tion of the LNG is an important factor in the solid formation concentration of this compound and other components in Fig. 16-39. Typically, reduction of the benzene concentration to 10 ppm is sufficient to prevent solid formation. The GPA performed research in this area (See Section 1) and has produced a predic-tive computer program to calculate freezing points for both hy-drocarbons and CO2 in LNG streams.

NITROGEN REJECTIONVirtually all natural gas contains some amount of nitrogen

whVirtually all natural gas contains some amount of nitrogen which lowers the heating value of the gas, but is so low to be no particular problem. However, in some reservoirs gas contains larger amounts of nitrogen than cannot be tolerated due to con-tractual considerations on minimum heating content. In these cases, the operator has three options:

1. Blend the gas with richer gas to maintain overall heating value;

2. Accept a reduced price for sales gas or a less secure mar-ket; or

3. Remove the nitrogen to meet sales specifications. Options 1 and 2 are reasonable approaches to the problem but are very location specific.

When a nitrogen rejection unit (NRU) is selected as a pro-cess option for a gas stream, it may be combined with NGL re-covery in an integrated plant design. A block flow diagram of a combined NGL/NRU facility is shown in Fig. 16-40.16 The over-all objective of this facility is to produce a nitrogen vent stream with minimum hydrocarbon content (that is normally sent to atmosphere), specification sales gas stream, and a specification NGL product. A primary contributor to facility cost is the re-quired gas compression. Regardless of the technology, recom-pression of the sales gas is usually required unless the residue gas can be marketed at 2050 kPa (ga) or less. Inlet compres-sion is necessary if the gas is available at much less than 2750 kPa (ga) and can be justified for higher pressure gas depending on the nitrogen rejection technology.

Cryogenic TechnologyNitrogen rejection is typically carried out using cryogenic

distillation technology to achieve very high hydrocarbon recov-ery and minimize methane losses in the nitrogen vent stream. Due to the low temperature operation the gas, after optionally being compressed to required inlet pressure, is fed to a pretreat-ment unit for CO2 and water removal. The CO2 will start freez-ing at –57°C, and therefore must be removed. For some technol-ogy the CO2 must be removed to 30 ppmv but modern designs remove part of the hydrocarbon product at warmer tempera-tures in a prefractionator and can tolerate higher CO2 levels. Typically, CO2 removal is accomplished with amine treating which can easily remove CO2 to acceptable levels. See Section 21 for details concerning CO2 removal. The dehydration step is carried out with molecular sieve dehydration, which is covered in Section 20. Another impurity requiring pretreatment to re-move is mercury, which attacks the aluminum heat exchangers

in the low temperature section. Mercury removal techniques are covered in Section 21. Typically, removal is accomplished with an adsorbent bed.

The process design of the nitrogen rejection unit is a strong function of the feed gas nitrogen content and should consider the sales gas compression power requirement to give an opti-mized design. For nitrogen contents below about 20%, a single column with a heat pump cycle has been used such as shown in Fig. 16-41. By operating the column at 2400 kPa (ga), low pres-sure methane can be used in the condenser. The drawback to this process is the heat pump compressor that is required to minimize overall power consumption. The heat pump system includes one or more intercondensers on the column resulting in increased plant complexity. At nitrogen contents above about 30% an interlinked two column system such as in Fig. 16-42 is the preferred choice. This design uses the nitrogen content of the feed gas as column reflux and is efficient due to the con-denser/reboiler arrangement. Lower feed gas nitrogen content leads to an increased loss of methane in the column overheads unless this basic process is modified. The high pressure (HP) column operates at about 2400 kPa (ga). This design is quite flexible and can be used at nitrogen contents above 50%. A re-cycle compressor can be added to handle nitrogen contents be-low 20% though this adds to power consumption. To reduce power consumption and processing cost for feed gas with low nitrogen content, prefractionator designs were developed whereby fractionation is performed in a first, relatively high pressure column to produce a low nitrogen content bottoms stream, with a heat pump system or an interlinked two column system downstream to effect almost total removal of hydrocar-

FIG. 16-39

Problem Compounds in LNG

Freezing Temp, °CCyclohexane 6.55

Benzene 5.53n-Decane –29.63n-Nonane –53.48n-Octane –56.76

n-Heptane –90.55Toluene –94.98

n-Hexane –95.31

FIG. 16-40

Nitrogen Rejection Flow Diagram14

InletGas

InletCompression

Sales Gas4825 kPa (ga)

Sales GasCompression

NGL

NGL / NRU

4825 kPa (ga) N2 Vent

DehydrationCO2Removal

CO2

16-34

bon from vented nitrogen. Newer licensed designs use a prefrac-tionator with a downstream column at intermediate pressure and use the evaporation of product hydrocarbon at one or more pressure levels to minimize overall compression duty.

Early NRU designs typically produced all the hydrocarbon product at 690 kPa (ga) or less whereas newer designs can pro-duce up to about half the hydrocarbon at feed gas pressure and the balance at intermediate pressures. These improvements have significantly reduced power requirements, capital cost and operating costs.

Variable nitrogen content feed gas requires careful design to ensure efficient operation over a range of compositions. In en-hanced oil recovery (EOR) applications, where the feed gas ni-trogen content can range from less than 5% to about 80%, an interlinked two column process may be optimal. However, the heat pump based process can produce nitrogen at 2400 kPa (ga), so a tradeoff study is usually required in EOR applications to decide which process is the optimum one to use.

The NGL recovery section may be designed for ethane and heavier recovery or propane and heavier recovery. Since NGL recovery is also a cryogenic process, it is possible to integrate it with the nitrogen rejection process to reduce capital and operat-ing cost compared to separate NGL extraction and NRU facili-ties. The NGL recovery is a traditional turboexpander setup except that the front-end heat exchange is integrated with the nitrogen and sales gas streams from the NRU section. The in-cremental cost for NGL recovery may be quite small, because many of the required process steps such as dehydration and compression are already present although the operation can be significantly more complex with multiple side streams feeding the sales gas recompressor.

As such integrated plants need to be carefully designed for variations in feed gas compositions to ensure they can meet both NGL product and sales gas specifications with minimal hydrocarbon content in the nitrogen stream.

Recovery EfficienciesIn the separation of nitrogen from natural gas, high purity

products are readily achievable. Sales gas purity of 2% or lower nitrogen is achievable. The hydrocarbon content of the nitrogen vent stream is typically specified at about 1% maximum to min-imize environmental issues. This is equivalent to over 99.95% hydrocarbon recovery for a low nitrogen content feed gas. NGL recovery efficiencies associated with an integrated NGL/NRU can be quite high as well. Ethane recovery of well over 90% with virtually complete propane and heavier recovery is typical. If ethane recovery is not desired, the process can be designed for high C3+ recovery and incidental ethane recovery.

Molecular SievesEngelhard developed the Molecular Gate® technology using

titanium silicate materials whose pore size can be adjusted to +/– 0.1 Angstrom. This technology is currently available for li-cense from Guild Associates. These materials can be used to preferentially adsorb molecules on a size exclusion basis. For example, a material with 3.7 Angstrom pore diameter will ad-sorb nitrogen with a molecular diameter of 3.6 Angstrom but pass methane with a molecular diameter of around 3.8 Ang-stroms. These units run best with an inlet pressure of 690 kPa (ga), but can operate from 480 to 900 kPa (ga).

FIG. 16-41

Single-Column NRU16

FIG. 16-42

Two-Column NRU16

16-35

The Molecular Gate® system utilizes a pressure swing adsorption (PSA) process. It can also be designed for the si-multaneous removal of CO2 and nitrogen from natural gases containing both impurities.17

ENHANCED OIL RECOVERYIn order to increase oil production in many reservoirs, the

injection of gas for enhanced oil recovery (EOR) has been car-ried out in numerous projects. The gas injection plan can lead to three different types of processing facilities. First, high meth-ane or high nitrogen gas can be injected for pressure mainte-nance of the reservoir. In this case the gas is in a separate phase from the oil phase, and any gas produced is simply recycled to the reservoir. Processing of the gas in traditional gas processing facilities is often carried out.

Second, the gas injected may be nitrogen with little or no hydrocarbons. In this case the injection conditions are chosen such that the nitrogen becomes miscible with the oil phase. As the oil is produced, the nitrogen and associated gas are pro-duced as a mixed gas phase. This produced gas can be reinjected or processed for fuel, sales gas and NGL production. The pro-cesses used separating the nitrogen in this case are described in the Nitrogen Rejection subsection previously covered.

The third type of EOR process involves the injection of CO2. Large volumes of CO2 are injected into the reservoir and be-come miscible with the oil phase. This CO2 essentially scrubs the oil from the reservoir and can greatly increase oil produc-tion. As with the miscible nitrogen injection projects, the CO2 is produced with the oil and gas and must be handled in the gas processing facilities. The CO2 that is injected into the reservoir

is typically purchased from third party suppliers and is the single greatest operating cost in the EOR project. Therefore, the CO2 produced with the associated gas is valuable and must be recovered and recycled to the reservoir.

CO2 Processing for EORThe CO2 produced in an EOR project can be separated from

the hydrocarbon components using solvent or membrane pro-cesses as described in Section 21 of this Data Book. However, solvent processes such as amines, potassium carbonate, and physical solvents, as well as membrane systems, were not de-signed to handle the large volumes of CO2, which are present in the EOR gas. The capital and operating costs of these systems increase in proportion to the acid gas content. Additionally, the CO2 is produced at low pressure and typically saturated with water. The EOR project needs high pressure, dry CO2 for rein-jection.

An EOR gas processing plant is designed for three primary separations. First, the methane in the gas is needed for fuel and possibly to sell for additional revenue. Second, the produced gas often contains hydrogen sulfide (H2S) which is removed from the CO2 stream for safety considerations. Third, EOR gas is typically rich in recoverable NGLs. Fig. 16-43 is an example EOR production profile. This example shows the effect of the EOR operations on the gas to be handled. The CO2 may start out at a few percent but eventually builds to over 90% as the gas volume increases. The NGL curve in Fig. 16-44 (on a CO2 free basis) shows that the hydrocarbon portion of the gas gets con-tinually richer. In fact, in most projects, the in-situ oil is actu-ally stripped of the midrange hydrocarbons such that over 10% of the crude production is in the gas phase with the CO2.

FIG. 16-43

Example EOR Production Forecast18

16-36

All of the required separations could be performed in a frac-tionation process which would produce dry CO2 at elevated pressure as one of the products. Each step of the separation of C1, CO2, H2S, and C2+ components has technology issues which must be addressed with non-traditional concepts to achieve the necessary separations by fractionation.

Separation of CO2 and MethaneThe relative volatility of CO2 and methane at typical operat-

ing pressures is quite high, usually about 5 to 1. From this standpoint, distillative separation should be quite easy. How-ever, at processing conditions, the CO2 will form a solid phase if the distillation is carried out to the point of producing high pu-rity methane. The phase equilibria considerations in this sepa-ration are discussed in detail in Section 25 of this Data Book.

Fig. 16-44 illustrates the theoretical limits of methane pu-rity, which can be obtained in a binary CO2/methane system. In practice the purity limits of the methane product are around 15% CO2.

One approach to solving this methane-CO2 distillation prob-lem is to use an extractive distillation approach known as the Ryan/Holmes process. This concept involves adding a heavier hydrocarbon stream to the condenser in a fractionation column. The addition of this stream, which can contain ethane and heavier hydrocarbons, significantly alters the solubility charac-teristics of the system such that virtually any purity of methane can be produced.

Fig. 16-45 illustrates the effect of adding a third component (in this case n-butane) to a CO2-methane distillation column producing 2% CO2 overhead. By adding n-butane, a column op-

eration profile without CO2 solid formation can be achieved. Adding greater amounts of the additive increases the safety margin away from the CO2 solid formation region. Other char-acteristics of this additive addition concept include:

• Raising the operating temperature of the overhead• Increasing CO2/methane relative volatility• Permitting higher pressure operation by raising the mix-

ture critical pressureIn practice the additive has been increased to the point that

propane refrigeration can be used for the overhead condenser rather than cascade refrigeration. Under these conditions the design acts much like a refrigerated lean oil process for NGL recovery.

CO2-Ethane SeparationThe separation of CO2 and ethane by distillation is limited

by the azeotrope formation between these components. An azeo-tropic composition of approximately 67% CO2, 33% ethane is formed at virtually any pressure.

Fig. 16-46 shows the CO2-ethane system at two different pressures. The binary is a minimum boiling azeotrope at both pressures with a composition of about two thirds CO2 and one third ethane. Thus, any attempt to separate CO2 and ethane to nearly pure components by distillation cannot be achieved by traditional methods. Extractive distillation is required.

As developed by Ryan and Holmes, the technique involves the addition of a heavier hydrocarbon, usually butane or heavi-er, to the top section of the distillation column.

FIG. 16-44

Distillation Profile CH4–CO2 Binary19

FIG. 16-45

Distillation Profile Binary Feed with nC4 Additive20

16-37

The upper dashed line in Fig. 16-46 represents the phase behavior of a multicomponent feed distilled with a butane-plus additive. With this technique, virtually any purity of CO2 and ethane is thermodynamically possible.

For the CO2-methane separation, the additive is introduced in the condenser. In the CO2-ethane separation, the additive is normally introduced several trays below the top of the column. The primary CO2-ethane distillation is achieved below the additive feed tray. It is in this area that the relative volatil-ity of the CO2 to ethane is reversed to remain above 1.0 and the azeotrope is circumvented. High relative volatilities are ob-tained at all points on and below the additive feed tray.

In the top portion of the column above the additive feed tray, no resolution of the azeotrope is achieved, as the relative vola-tility of CO2/ethane is less than 1.0. This part of the column serves as a recovery zone for the extractive distillation addi-tive.

While the separation of these species is feasible as described, the value of the ethane may not justify extracting it from the recycled CO2.

Separation of CO2 and H2SThe distillative separation of CO2 and H2S can be performed

with traditional methods. The relative volatility of CO2 and H2S is quite small. While an azeotrope between H2S and CO2 does not exist, the vapor liquid equilibrium behavior for this binary approaches azeotropic character at high CO2 concentrations.21

In many cases the CO2 is required to contain less than 100 ppmv H2S. In order to achieve such purity a very large fraction-ation tower is required with large energy requirements.

Another aspect to be considered is the CO2 in the bottom (H2S concentrated) stream. If fed to a Claus sulfur recovery plant, the CO2/H2S ratio is desired to be less than 2 to 1. Achiev-ing such a low ratio will require high energy input in many cases.

By adding a third component, as in the CO2-ethane separa-tion system, the relative volatility of CO2 to H2S is significantly enhanced. Fig. 16-47 demonstrates the relative volatility en-hancement due to addition of n-butane to the CO2–H2S binary system.

Thus, if a system containing CO2, ethane and H2S is pro-cessed in an extractive distillation column, the ethane and H2S can be separated from the CO2. The exact specification for pu-rity and recovery will determine the system operating require-ments. From a thermodynamic standpoint the CO2 could be produced overhead with the ethane and the H2S (and any C3+ components) produced as a mixed bottom product.

FIG. 16-46

Vapor-Liquid Equilibria CO2–C2H620

FIG. 16-47

CO2–H2S–nC4 System at 4100 kPa22

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Again this separation may not be economically attractive, especially since there are H2S selective solvents available for extracting H2S, while leaving the bulk of the CO2. If this route is elected, then this process should be done on the raw gas prior to dehydration, since these solvents all contain water as a con-stituent. See Section 21 for more details.

Overall Process ConfigurationThe EOR processing steps can be arranged in a system to

achieve all the desired separations although the process con-figuration can take on several variations, the configuration most often used in EOR processing plant is shown in Fig. 16-48.

In this configuration, the first step is the ethane/CO2 sepa-ration in the ethane recovery column. This separation is carried out at pressures in the 2400 kPa (ga) range using refrigeration for reflux in the –18°C range. The CO2 and lighter components are taken overhead, compressed to around 4500 kPa (ga) and sent to the CO2 recovery column. This is a bulk removal column which produces CO2 as a liquid bottom product. This CO2 can then be pumped to reinjection. The overhead product is essen-tially a CO2/C1 binary which is limited by CO2 solid formation considerations. This binary stream is then separated by use of the extractive distillation step to produce a methane stream with low CO2 content.

The bottoms products from the ethane recovery and demeth-anizer columns are combined and processed in the additive re-

covery column. In this column the additive, which is a C4+ stream, is separated from the lighter hydrocarbons for recycle to the distillation columns. A net C4+ product is also produced. The additive used in the ethane recovery and demethanizer col-umns is continuously regenerated and reused much the same as lean oil in traditional gas processing applications. The dis-tinct difference in this case is that this C4+ stream is generated from the feed gas and is used as an extractive distillation agent rather than as an absorption agent.

The light NGL product produced overhead in the additive recovery column also contains any H2S which was present in the feed gas and some residual CO2. This product is usually treated in a small amine or dry bed unit to meet sales specifica-tions. The acid gas may then be sent to a sulfur recovery unit.

This four column EOR processing plant is designed to han-dle the wide range of feed rates and compositions encountered in EOR projects. In the design effort early, peak and late year cases must be investigated to ensure proper operation over time. This is especially important since the exact timing, flow rate and composition of EOR production are extremely difficult to predict. CO2 breakthrough to the processing plant can occur rapidly. In some projects the CO2 volume can triple in less than one year. As the EOR process has matured, other configura-tions have been developed which mix technologies such as membranes with Ryan/Holmes facilities to optimize the capital and operating costs over the project life.

FIG. 16-48

Four-Column Ryan/Holmes Process18

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REFERENCES 1. Rusten, B.H., Gjertsen, L.H., Solbraa E., Kirkerød T., Toril Hau-

gum T. and Puntervold S., “Determination Of The Phase Enve-lope – Crucial For Process Design And Problem Solving,” 87th GPA Annual Convention, Grapevine, TX , March 2-5, 2008.

2. Maddox, R.N., and Erbar, J.H., ”Low-Pressure Retrograde Condensa-tion,” Oil & Gas Journal, July 1977.

3. Petroleum Extension Service, “Field Handling of Natural Gas,” Third Edition, University of Texas, Austin, TX, 1972, pp. 36-51.

4. Sachin, J. and Lokhandwala, K., “Heavy or Sour Fuel Gas? Max-imize Uptime Using a Simple Membrane Process to Improve Gas Quality,” 2010 Gas Machinery Conference, Phoenix, AZ, October 4-6, 2010.

5. Toreja, J., VanNostrand, B., Chan, N. and Dickinson, J.P., “Ro-tary-valve, Fast-cycle Pressure-swing Adsorption Technology Al-lows West Coast Platform to Meet Tight California Specifications and Recover Stranded Gas,” 61st Laurence Reid Gas Condition-ing Conference, Norman, OK, February 2011.

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11. Nasir, P., Sweet, W., Elliot, D., Chen, R. and Lee, R.J., “Enhanced NGL Recovery Processsm Selected for Neptune Gas Plant Expan-sion,” 82nd GPA Annual Convention, March 9-12, 2003.

12. Ulmann’s Encyclopedia of Industrial Chemistry, VCH Publish-ers, New York, NY, 1991, Vol 17, pp. 100-109.

13. Price, B.C. and Mortko, R.A., “PRICO – A Simple, Flexible Proven Approach to Natural Gas Liquefaction,” Gastech 96, Vienna, Aus-tria, December 3-6, 1996.

14. Chatterjee, N, Kinard, G.E., and Geist, J.M., “Maximizing Pro-duction in Propane Precooled Mixed Refrigerant LNG Plants,” Seventh Conference on Liquefied Natural Gas, Jakarta, Indonesia, May 15-19, 1983.

15. Finn, A. J., “Floating LNG Plants – Scaleup of Familiar Technol-ogy”, 88th GPA Annual Convention, San Antonio, TX, March 2009.

16. Price, B.C., “NRU’s Upgrade Production, Cut Costs,” American Oil and Gas Reporter, March 1994, pp. 56-60.

17. Mitariten, M.,Dolan, W. , and Maglio, A., “Innovative Molecular Gate Systems for Nitrogen Rejection Carbon Dioxide Removal and NGL Recovery,” 81st GPA Annual Convention, Dallas, Texas, March 2002.

18. Price, B.C., “Looking at CO2 Recovery in Enhanced Oil Recovery Projects,” Oil & Gas Journal, December 24, 1984, pp. 48-53.

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20. Holmes, A.S., Ryan, J.M., Price, B.C., and Styring, R.E., “Pilot Tests Prove Out Cryogenic Acid-Gas/Hydrocarbon Separation Processes,” 61st Annual GPA Convention, Dallas, TX, March 15-17, 1982.

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22. Ryan, J.M., and Holmes, A. S., “ Distillation Separation of Carbon Di-oxide from Hydrogen Sulfide,” U.S. Patent No. 4,383,841 (1983).

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NOTES: