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New NGL-recovery processprovides viable alternativeRobert R. HuebelMichael G. MalsamRandall Gas TechnologiesHouston

can achieve NGL recovery efficiencies comparable to that of advanced turbo-expander cycles but for lower capital and operating expenditures.

This article reviews the fundamen-tals of the IPOR process, including process features, benefits, and appli-cability. It also presents case studies that compare process performance with both straight refrigeration and advanced turboexpander cycles and economic analysis.

Diverse environmentsNatural gas conditioning and process-ing plants are somewhat unique in that the raw material feedstock is typically fed into the plant at the pressure, flow rate, and composition at which it is produced.

Consequently, natural gas process-ing plants have considerable variation in size, complexity, and configuration, depending upon specific reservoir production characteristics, geography, customer specifications, and market

drivers. These range from simple dewpoint plants with ca-pacities less than 5 MMscfd and minimal hydrocarbon re-covery to large deep cut ethane extraction straddle plants which process in excess of 1 bcfd .

With such a diverse operating environment, it is a bit sur-

Operational scenarios for two uses of a new refrigeration process for recover-ing NGLs from natural gas have shown it to enhance operability and reduce capital and operating expenditures when compared with the two more traditional process choicesstraight refrigeration and turboexpander.

Straight refrigeration units that most often use propane as refrigerant have proven to be economical and re-liable. Their operating temperature, however, typically about 35 F., limits NGL extraction. For higher NGL re-covery, todays processor is left with a cryogenic turboexpander.

IPOR (IsoPressure Open Refrigera-tion) has been developed by Randall Gas Technologies, a division of Lum-mus Technology, a CB&I company, to bridge this gap. The advanced re-frigeration process can economically achieve essentially total C

3+ recovery

from most natural gas streams. Using conventional closed-loop mechani-cal refrigeration combined with an open-loop mixed refrigeration cycle, the new technology

Based on a presentation to the GPA Europe Annual Conference, Prague, Sept. 21-23, 2011.

CASE STUDY 1:PLANT DESIGN BASIS

Table 1

Feed gas: Flow, MMscfd 100 Pressure, psig 70 Temperature, oF. 80 Composition, mol %: N2 & CO2 1.8 C1 75.0 C2 16.1 C3+ 7.3Residue gas: Pressure, psig 1,200 Heating value, btu/scf max 1,100NGL product specifications: C2/C3 liquid volume ratio 0.02

CASE STUDY 2:PLANT DESIGN BASIS

Table 2

Feed gas: Flow, MMscfd 20 Pressure, psig 200 Temperature, oF. 50 Composition, mol %: N2 & CO2 1.9 C1 81.2 C2 9.3 C3+ 7.6Residue gas: Pressure, psig 950 Hydrocarbon dewpoint, oF. 5NGL product specifications: C2/C3 liquid volume ratio 0.02

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prising that natural gas processors have had essentially only two process technology choices for extracting hydrocarbon liquids from natural gas: either straight refrigeration or tur-boexpander. Among more than 1,600 operating natural gas processing plants shown in Oil & Gas Journals Worldwide Gas Plant Survey, about 80% use either straight refrigeration or turboexpander technology (OGJ, June 6, 2011, p. 88).

With the last new lean oil plant built some 30 years ago, the estimated portion of new gas plants built today using these two technologies is greater than 95%.

Straight refrigeration units, which most often use pro-pane or ammonia, can be built for essentially any capacity or feed-gas composition, are of mild steel construction, are relatively simple to construct and operate, and have proven to be economical and reliable. However, with their operating temperature typically limited to about 35 F., their capabil-ity for NGL extraction is limited.

For higher NGL recovery, todays processor has but a single choice: cryogenic turboexpander. Since its inception in the late 1960s, turboexpander technology has evolved into the technology of choice for deep NGL-product recov-ery. As designs were refined, turboexpander technology essentially displaced lean-oil technology for high LPG or ethane-extraction applications.

Several variations of the technology are available, depend-ing upon the targeted product recovery and feed-gas condi-tions, with proprietary designs offering even higher efficien-cies. With operating temperatures as low as 200 F., NGL product recoveries approaching 98%+ are technically feasible.

With straight refrigeration technology, the benefits for the customer include low capital and operating expenditures (CAPEX and OPEX), a broad range of applicability, early production capabilities, but limited NGL recovery. Expander technology offers superior NGL-recovery potential but high-

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FIG. 1IPOR PROCESS FOR HIGH LPG RECOVERY

Gas-gasexchanger

NGL

SalesFeed Mixed refrigerantcompressor

Propane refrigerationcompressor

De-ethanizerreux drum

De-ethanizerreboiler

De-ethanizeroverheadseparator

De-ethanizer

De-ethanizeroverheadcondenser

Mixed refrigerantgas-gas exchanger

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Feed gas, at a pressure typically 300-550 psig, is initially cooled and partially condensed in the gas-gas heat exchang-er by cross exchange with cold residue gas and propane re-frigerant. A conventional brazed aluminum heat exchanger appears in the flow diagram; however, shell-and-tube ex-changers can also be used for this service.

The cooled and partially condensed feed-gas stream is then fed to the middle section of the de-ethanizer, which uses either trays or packing or a combination of these to effect the desired product separation. Below the feed tray, the stripping section of the column selectively removes the lighter fractions to meet product specifications, which nor-mally is 2-5% ethane in the recovered propane. Heat for the separation is provided by the de-ethanizer reboiler, which is a conventional shell-and-tube heat exchanger, with the heat supplied from the plant heating medium system.

In the upper section of the de-ethanizer, above the feed tray, the cooled feed gas flows counter-currently to the reflux stream, which is fed to the top tray in a conventional man-ner. The reflux provides additional cooling for the feed-gas stream and also selectively absorbs the propane and heavier components from the gas, thereby providing high product recovery efficiencies.

The overhead gas stream from the de-ethanizer, at this point in the process containing primarily the light ends from the feed-gas stream and a small portion of the propane, is

er CAPEX and OPEX and a longer time to initial operation due to the long lead time of such specialty equipment as the turboexpander and brazed aluminum heat exchangers.

Ethane-rich cycleThe advanced refrigeration NGL extraction process can eco-nomically achieve deep NGL extraction from most natural gas streams. Using conventional closed-loop mechanical re-frigeration combined with an open-loop mixed refrigeration cycle, this process can provide performance comparable to that of advanced turboexpander technologies but with much lower CAPEX and OPEX.

Unique about the IPOR process is its open-loop ethane-rich mixed refrigeration cycle. This refrigerant, extracted from the feed gas itself, is a mixture of predominantly ethane with lower concentrations of methane, propane, and other feed-gas constituents.

This refrigeration cycle serves a dual purpose: producing the cryogenic refrigeration for the process to enable lower tempera-ture operation while at the same time providing a reflux stream to the fractionation column, the combination of which produc-es high product extraction and thermal efficiencies.

The extraction process can be configured in several ways, depending on the feed stream, site conditions, and project objectives. Fig. 1 depicts one configuration of the IPOR tech-nology recommended for high recovery LPG applications.

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FIG. 2TURBOEXPANDER PROCESS

129 F.

NGL

Fuel

Residue gas torecompression

(160 psig)

Feed (400 psig)

Residuerecycle

Reboilers

Coldseparator

Boostercompressor-

expander

De-ethanizerreboiler

De-ethanizerD

emethanizer

Propanerefrigerant

35 F.

58 F.

PROCESSING

requirements in the de-ethanizer overhead condenser and to minimize the compression power requirements.

From the de-ethanizer overhead condenser, the mixed refrigerant stream is heated further as it flows through the mixed refrigerant gas-gas exchanger to the mixed refriger-ant compressor. The discharge pressure of this compressor is normally about 40 psig higher than the operating pressure of the de-ethanizer.

The mixed refrigerant compressor is of conventional design and can be either reciprocating, centrifugal, or screw type, de-pending upon project requirements and customer preferenc-es. Drivers may be gas turbine, gas engine, or electric motor. The compressor can be packaged with driver, scrubbers, and discharge cooler following standard industry practice.

The compressed ethane-rich, mixed refrigerant stream is then cooled and partially condensed in the mixed refriger-ant gas-gas exchanger. Cooling for this exchanger is provid-

further cooled in the de-ethanizer overhead condenser by cross exchange with cold residue gas and the ethane-rich mixed refrigerant stream.

The cooled and partially condensed gas stream flows to the de-ethanizer overhead separator. The liquid from this separation, a mixture of methane, ethane, and propane, is used as the refrigerant for the open-loop mixed refrigerant cycle. The de-ethanizer overhead separator therefore has a twofold function: It acts as a conventional two-phase gas-liquid separator, and it provides surge capacity for the liquid mixed refrigerant system.

From the de-ethanizer overhead separator, the pressure of the liquid mixed refrigerant is reduced, creating a Joule-Thomson refrigeration effect: This cold stream provides the desired cooling in the de-ethanizer overhead condenser. The pressure of the low-pressure mixed refrigerant, usually in the range of 100-200 psig, is selected to satisfy the cooling

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FIG. 3IPOR PROCESS

Gas-gasexchanger

NGL

SalesFeed

Fuel

Mixed refrigerantcompressor

Propane refrigerantcompressor

De-ethanizerreux drum

De-ethanizerreboiler

De-ethanizeroverheadseparator

De-ethanizer

365 psig

410 psig

120 psig

42 F.

10 F.

75 F.

De-ethanizeroverheadcondenser

Mixed refrigerantgas-gas exchanger

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IPOR process, the process refrigeration temperature is in the range of 10 F. to 20 F.; other refrigerants, therefore, such as ammonia may be used as well.

For the LPG-recovery configuration above, product extrac-tion efficiencies are excellent, with C

3 recovery in the range of

95-99%+, with essentially 100% recovery of the C4+ fraction.

From a thermal efficiency perspective, the IPOR process requires about 15-40% less compression power than a com-parable turboexpander design. As a result, plants using the IPOR technology will also have lower emissions and a small-er carbon footprint.

ed by low-temperature mixed refrigerant and propane. The two-phase stream flows to the de-ethanizer reflux drum, a conventional two-phase gas liquid separator. This liquid is used to provide reflux to the de-ethanizer column, thereby completing the open cycle of the mixed refrigerant loop.

Noncondensable vapors, consisting mainly of methane, are directed back into the process via the de-ethanizer over-head separator and eventually exit the process into the resi-due gas stream or may be used as fuel.

The closed-loop propane refrigeration is of conventional natural gas industry design and construction. In a typical

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FIG. 4MECHANICAL REFRIGERATION

C3+ product

Reboiler

Coldseparator

Cooler

Recompressor

Gas-refrigerantexchanger

Gas-gasexchanger

Residue gas

Fee gas (975 psig)

10 F.

225 psig

Gas chiller

Stabilizer

CASE STUDY 1: RESULTS Table 3 IPOR Turbo-Feed-gas capacity: 100 MMscfd process expander

Product recovery: C3 99.5 98.8 C4+ 100.0 100.0 NGL production, b/d 5,173 5,131

Power, bhp: Inlet compression 9,780 10,460 Residue compression 7,060 11,930 Refrigeration 4,730 1,910 Pumps, air coolers 490 560 Total power 22,020 24,860

Gas compression 16,000 16,000 Process compression 6,020 8,860

Major equipment countprocess: Turboexpander 1 Pumps 4 Columns 1 2 All other 24 24 Total major equipment countprocess 25 31

CASE STUDY 1: ECONOMIC ANALYSIS Table 4 IPOR Turbo-Feed-gas capacity: 100 MMscfd process, $ expander

CAPEX

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ited only by the performance of in-line control instruments, i.e., control valves, meters, etc., unlike turboexpander designs, which suffer from an inherent loss of efficiency at reduced flows.

The process can be designed for a wide variety of feed-gas compositions, site conditions, and capacities. Ethane recov-ery can be incorporated into an IPOR process design, with ethane recoveries up to 80%, depending upon feed-gas com-position. Equipment can be incorporated to allow for future ethane recovery, or the initial design can permit operation in ethane-rejection/ethane-recovery mode.

The process was developed based on proven technolo-gies and equipment employed extensively in gas plants. All the equipment incorporated into the process design is well within the natural gas industrys experience and capability. The low equipment count, small footprint, and process sim-plicity of the technology permit a compact layout and a high degree of modularization.

The process utilizes equipment and materials that are all well proven within the natural gas processing industry. Most of the unit can be of carbon steel or low-temperature carbon steel construction; typically the only major equipment item that requires stainless steel construction is the de-ethanizer overhead separator.

The only rotating equipment required for the IPOR process is the refrigerant compressor. The process requires no cryo-genic turboexpander or light hydrocarbon pumps. As a result:

Reliabilityandoperabilitywillbecomparable to thatof a conventional refrigeration process and should exceed that of a modern day turboexpander facility, given the fewer items of rotating equipment.

Theprocessofferssuperioreconomicsforalmostanyfeed-gas rate, from as low as 5 MMscfd to 1 bcfd+.

Almost infinite turndown capacity is possible with anIPOR process, to feed-gas rates as low as 10% of design, lim-

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FIG. 5IPOR PROCESS FOR CASE STUDY 2

Gas-gasexchanger

Residue gasto recompressor

NGL

FeedMixed refrigerant

compressorPropane refrigerant

compressor

De-ethanizerreux drum

De-ethanizerreboiler

De-ethanizeroverheadseparator

De-ethanizer

De-ethanizeroverheadcondenser

Mixed refrigerantgas-gas exchanger

Gas chillerGas chiller

410 psig

20 F.

125 psig

435 psig

105 F.

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anizer column is consumed as fuel to achieve the residue-gas heating value specifications.

Fig. 3 illustrates the IPOR process used in the study. Feed gas enters the process unit at a compressor interstage pres-sure of about 365 psig. The propane refrigeration system op-erates at 10 F., much warmer than that required by the tur-boexpander process, and 16.7 psig. The minimum operating temperature of the de-ethanizer column is 42 F. and is of low-temperature carbon steel construction.

To achieve the residue-gas pipeline heating value specifi-cation, a portion of the ethane-rich noncondensable vapors from the de-ethanizer reflux drum is consumed as fuel, with the remainder mixing with the residue gas via the de-etha-nizer overhead separator.

Tables 3 and 4 summarize the results of the study.Compared with the turboexpander design, the IPOR

Marcellus plantA recent study compared the IPOR process with modern tur-boexpander technology. Feedstock for the new plant is from the Marcellus shale, a region with limited existing oil and gas infrastructure and no existing ethane market. Demand for LPG in the region is strong, with extracted LPG sold into the local market.

As a result, the customer wanted to maximize LPG pro-duction. Due to the richness of the gas, some ethane extrac-tion was required to meet the residue-gas pipeline specifica-tions, with the ethane consumed within the plant as fuel. The fields gathering system operated at low pressure, with residue gas delivered into an existing high pressure pipeline.

Table 1 summarizes the design basis for the plant. Two process technologies were evaluated: conventional

turboexpander and the IPOR process. The turboexpander process utilized in the study was a

modern design (Fig. 2). Due to the richness of the feed gas, a propane refrigeration system with a low stage operating temperature of 35 F. at 3.4 psig was integrated into the process design to provide supplemental cooling. A portion of the ethane vapor stream from the overhead of the de-eth-

PROCESS COMPARISONS Table 7 Refrigeration IPOR process Turboexpander

Applicability Feed-gas volume, MMscfd Any Any 50+ Feed-gas pressure, psig Any 2.0 million Base Mole-sieve dehy vs. glycol Additional plate fins Additional compression OPEX (fuel) >150,000/year Base Fuel value @ $4.50/MMbtu NGL revenue >5.1 million/year Base Shrinkage @ $4.50/MMbtu Crude @ $80/bbl C3 @ 60% of crude C4 @ 80% of crude C5 @ 90% of crude Trans. & frac. @ $0.05/gal 96% availability Internal rate of return 155% Base 38% tax rate Double-declining balance depreciation rate 20-year plant life Payback period

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the feed-gas compression, with residue gas sent directly to the sales gas pipeline.

Fig. 5 illustrates the IPOR process used in the study. Feed gas enters the IPOR unit at a compressor interstage pressure of about 410 psig. For this design, the gas-gas exchangers and chillers were conventional shell-and-tube design. All of the noncondensable vapors from the de-ethanizer reflux drum flow to the de-ethanizer overhead separator and on to the residue gas stream.

Tables 5 and 6 summarize the studys results.From these results, key observations include the following:1. NGL production with the IPOR unit is more than dou-

ble that of the refrigeration plant.2. Complexity of the two designs is comparable,

process:1. Achieves higher NGL recovery. 2. Requires about 32% less process compression power.3. Requires about 20% less major equipment.4. Requires less rotating equipment.As a result, economics of the IPOR process are clearly su-

perior to the turboexpander design, both from an OPEX and a CAPEX perspective (Table 4). Estimated capital cost of the IPOR process design was $11 million less than that of the turboexpander plant, the savings being the result of:

1. Less installed compression.2. No turboexpander.3. No light hydrocarbon/cryogenic pumps.4. No stainless steel demethanizer column.5. Less alloy material.From an operating cost perspec-

tive, the IPOR process was esti-mated to consume about $700,000/year less in utilities, the savings resulting from lower compression power requirements, and hence fuel gas consumption.

Northwest CanadaA second study was recently com-pleted comparing the IPOR process to a straight refrigeration process. Location for this plant is in north-west Canada, an area of existing oil and gas production but no NGL or ethane pipeline infrastructure. Liq-uids produced in the plant would be trucked to market.

The primary objective of the cus-tomer in this application was to de-liver a marketable sales gas. Given the current favorable economic cli-mate for gas liquids, however, incre-mental LPG recovery was of interest if economical.

The basis of design of the plant for the study is discussed below. Ta-ble 2 summarizes the design basis.

The straight refrigeration pro-cess used in the study was a tradi-tional design (Fig. 4), with process temperature selected to achieve the pipeline dewpoint specification, thereby minimizing both CAPEX and OPEX. Propane was used as the refrigerant, with glycol injection used for hydrate inhibition and de-hydration. Feed gas for the refrigera-tion unit was taken downstream of

NELSON-fARRAR COST INDExESRefinery construction (1946 basis)

(Explained in OGJ, Dec. 30, 1985, p. 145, and at www.pennenergy.com/index/research-and_data/oil-and_gas/Statistic-Definitions.html; click Nelson-Farrar Cost Indices)

Sept. Aug. Sept. 1962 1980 2008 2009 2010 2010 2011 2011

Pumps, compressors, etc.222.5 777.3 1,949.8 2,011.4 2,030.7 2,036.4 2,119.6 2,120.5

Electrical machinery189.5 394.7 515.6 515.5 513.9 513.7 515.0 514.1

Internal-comb. engines183.4 512.6 990.9 1,023.0 1,027.8 1,021.2 1,036.3 1,036.3

Instruments214.8 587.3 1,342.1 1,394.8 1,435.1 1,437.3 1,458.2 1,461.6

Heat exchangers183.6 618.7 1,354.6 1,253.8 1,116.0 1,103.5 1,103.5 1,253.8

Misc. equip. average198.8 578.1 1,230.6 1,239.7 1,224.7 1,222.4 1,246.5 1,277.3

Materials component205.9 629.2 1,572.0 1,324.8 1,480.1 1,489.4 1,619.7 1,627.6

Labor component258.8 951.9 2,704.3 2,813.0 2,909.3 2,923.3 2,992.9 3,000.2

Refinery (Inflation) Index237.6 822.8 2,251.4 2,217.7 2,337.6 2,349.8 2,443.6 2,451.2

Refinery operating (1956 basis)(Explained in OGJ, Dec. 30, 1985, p. 145, and at www.pennenergy.com/index/research-and_data/oil-and_gas/Statistic-Definitions.html; click Nelson-Farrar Cost Indices)

Sept. Aug. Sept. 1962 1980 2008 2009 2010 2010 2011 2011

Fuel cost100.9 810.5 1,951.3 978.5 1,184.9 1,048.8 1,267.6 1,196.5

Labor cost93.9 200.5 237.9 264.5 281.7 277.9 255.1 262.4

Wages123.9 439.9 1,092.2 1,177.1 1,279.4 1,289.0 1,270.6 1,267.5

Productivity131.8 226.3 460.8 445.2 454.5 463.9 498.2 483.0

Invest., maint., etc.121.7 324.8 830.8 812.4 850 854.5 888.6 891.3

Chemical costs96.7 229.2 472.5 406.2 449.8 444.0 557.9 560.1

Operating indexes Refinery

103.7 312.7 674.1 582.6 628.2 615.9 652.2 650.0Process units*

103.6 457.5 1,045.1 706.1 796.8 749.5 831.2 809.5

*Add separate index(es) for chemicals, if any are used. See current Quarterly Costimating in first issues for January, April, July, and October. These indexes are published in the first of each month. They are compiled by Gary Farrar, OGJ Contributing Editor. Indexes of selected individual items of equipment and materials are also published on the Costimating page in first issues for January, April, July, and October.

PROCESSING

based upon major equipment count, which should re-sult in similar operability and reliability. (Major equip-ment count in this case includes the entire plant fa-cility, including dehydration, utilities, and off sites.)

As a result, economics of the IPOR process are once again superior to the refrigeration unit, taking into account incre-mental differences in both OPEX and CAPEX (Table 6). Esti-mated capital cost of the IPOR process design was $2 million more than that of the refrigeration plant, the additional cost being the result of:

1. More installed compression.2. Additional heat-exchanger costs.3. Additional cost of the molecular-sieve dehydration

system vs. the glycol injection system utilized in the refrig-eration plant design.

4. More alloy materials.The operating cost of the IPOR process was estimated to

be about $150,000/year more than the refrigeration process. The additional cost was primarily the result of the higher compression power requirements of the IPOR process, and therefore more fuel-gas consumption.

NGL production with the IPOR unit is more than double that of the refrigeration plant. The value of this addition-al NGL revenue was estimated at $5.1 million/year. Based upon the economic assumptions itemized in Table 6, the cal-culated internal rate of return of the IPOR plant investment is 155%, with a payback of fewer than 6 months.

While the IPOR unit requires somewhat more CAPEX and OPEX than a minimal type investment of the refrig-eration unit, these costs are more than compensated for with the increased NGL revenue.

Table 7 summarizes the comparisons discussed in this article.

The authorsRobert R. Huebel (rhuebel@cbi.com) is vice-president of technology of Randall Gas Technologies, a division of Lummus Technol-ogy Inc., a CB&I company. His previous role was as president of the ABB Randall Corp. He joined the company in 1976. Huebel has more than 40 years experience in the domestic and international engineering, pro-curement, and construction and natural gas processing industries, including process engineering, project management, contract development, and executive manage-ment. He holds a BSc in chemical engineering and an MBA from the University of Houston. Huebel is a registered profes-sional engineer in seven states, is a member of the American Institute of Chemical Engineers and Project Management Institute, and currently serves on the board of directors of the Gas Processors Suppliers Association.

Michael G. Malsam (mmalsam@cbi.com) is senior principal process engineering specialist for Randall Gas Technologies, which he joined in 1998. He has more than 30 years experi-ence in the domestic and international EPC and natural gas processing, including process engineering, project development, and project management. Malsam holds a BSc in chemi-cal and petroleum refining engineering from the

Colorado School of Mines. He is a member of the American Insti-tute of Chemical Engineers and the American Chemical Society.

References1. Malsam, Michael G, IPOR TechnologyA new means

of LPG recovery, Gas Processors Association Annual Con-vention, High Definition at 90Advancing the Midstream Vision, March 2011.

2. Gas Processing with Cryogenic Turboexpander Technol-ogy, Randall Gas Technologies, Houston; January 2011 Edition.

Reprinted with revisions to format, from the January 9, 2012 edition of Oil & Gas JournalCopyright 2012 by PennWell Corporation