Commercial applications for gas permeation membrane systems

Journal of Membrane Science, 22 (1985) 2 1 7-224

217

Elsevier Science Publishers B.V., Amsterdam - Printed in The Netherlands

COMMERCIAL APPLICATIONS FOR GAS PERMEATION MEMBRANE SYSTEMS*

W.J SCHELL

Separex Corporation, 2100 E Orangethorpe Avenue, Anaheim, CA 92806 (U .S .A ) (Received January 2, 1984, accepted February 27, 1984)

Summary The commercialization of membrane systems for gas separation has become a reality in recent years The asymmetric structure of these membranes provides high flux and selectivity in the same manner as has been realized in reverse osmosis . While similar principles apply in both processes, the parameters affecting separation are different . In reverse osmosis one is concerned with concentration and osmotic pressure, whereas in gas separation one is concerned with partial pressures and the pressure ratio of the feed gas to the permeate gas . As a process, membrane gas separation competes with adsorption, absorption and distillation technologies . In some cases, however, a combination of membranes with these processes offers economic or technical advantages [1,2] . In the following paragraphs several applications of membranes will be discussed in comparison with these conventional methods . In addition, a comparison of a cellulose acetate membrane and a dimethyl silicone membrane for nitrogen enrichment will be presented to illustrate how membrane performance affects economics .

Landfill gas recovery A typical landfill gas recovery scheme is shown in Fig . 1 . Gas produced from landfills typically contains 40-45 mol% CO 2 , 54-59 mol% methane and 0 .5-1 .0 mol% trace contaminants such as water, heavy hydrocarbons, aromatics and chlorinated hydrocarbons . The raw gas is collected under vacuum, transfered as much as two miles, compressed, treated for contaminant removal and separated into CO 2 and methane . The most suitable conventional technology currently available for the CO 2 separation step is the Selexol process, where the CO 2 is absorbed by the Selexol solvent at high pressure, and desorbed by reducing the pressure . A comparison of membranes with the Selexol process is shown in Table 1 . The numbers shown are normalized to a value of 1 for the membrane system . The feed in both cases is compressed from 7 psia to 400 psia and is reduced *Paper presented at the 4th Symposium on Synthetic Membranes in Science and Industry, Tiibingen, F .R .G ., September 6-9, 1983 .

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© 1985 Elsevier Science Publishers B .V .



218 in carbon dioxide content from 42% to 3 .5% . The costs shown do not include compressors or compressor power consumption, as these are identical for both cases . It can be seen from Table 1 that the membrane system is significantly lower both in installed capital cost and in operating cost . Although more methane is recovered with the Selexol process, the cost to produce a unit volume of methane is almost twice as high compared to the membrane system, CO 2 VENT GAS

COLLECTION UNDER VACUUM

,N COMPRESSION



SEPARATION

Heavy Hydrocarbons Contaminants water

LANDFILL

T I

1

SALES GAS

Fig . 1 . Schematic diagram for landfill gas recovery. TABLE 1 Cost comparison for landfill recovery

Installed capital cost Annual operating costs Membrane repl . Utilities Volume CH, produced Cost to produce CH 4

Membrane system

Selexol

1.00

1 .99

1 .00 1 .00 1 .00

2.70 1,16 1 .88

CO 2 removal from fractured wells Another membrane system application for the purification of a gas stream containing carbon dioxide is shown in Fig . 2. This case refers to the treatment of a gas produced from a fractured well, where a tight gas deposit is opened up by injecting high-pressure carbon dioxide and water into the well to fracture the formation [3] . From 500 to 1,200 tons of carbon dioxide is used in a typical well fracture, which increases gas production 10-20 fold [4] . When the well is first opened, the produced gas is highly enriched in carbon dioxide, and therefore unsuitable for sale to the pipeline . In order to reduce the carbon dioxide level to pipeline specification a membrane system may be temporarily employed to remove CO 2 until the well returns to pipeline quality. The graph shown in Fig . 2 gives an approximate composition of the feed gas as a function of time, and as a function of concentration of the residual



219

gas produced from the membrane system . Typical feed conditions are 900 to 1,500 psia (63 to 105 bar) at 25 .45 ° C . The produced feed gas is watersaturated and fairly constant in flow rate . The permeate gas containing CO 2 and removed water vapor is vented to the atmosphere at approximately 15 psia 50 .

FEED CONDITIONS 900-1500 PSIG 80.110°F WATER SATURATED VOLUME : CONSTANT MEMBRANE AREA' FIXED PERMEATE AT 15 PSIA

30, 20. 0 W W Ll-

10,

0

5

z n 0 0

3. 2. %C02 IN RESIDUAL 5 0

1 1

1 2

30

•

3

20

l 5

10

05

03

02 .

•

10 TIME, DAYS

20

30 50

100

Fig . 2 . Gas compositions from fractured well .

Fig . 3. Trailor-mounted Separex TM membrane system for removing CO 2 from natural gas .

220 (1 bar) . It can be seen that the gas stream from the well approaches a pipeline specification of 2% CO 2 after about 60 days of operation, and the membrane system can then be moved to another location for additional fracture treatment service . A trailer-mounted Separex T M system for this application that treats 3 .5 MMSCFD (173 N-m 3 /day) of gas is shown in Fig . 3 . A suitable conventional treating process for this application is a chemical solvent system utilizing diethanolamine (DEA) . A relative cost comparison of DEA with a membrane system for a 3 .5 MMSCFD feed stream is shown in Table 2. In this case methane recovery by the DEA process is somewhat higher than with the membrane process but the capital cost is almost three times as much . In addition, the DEA system is a permanent installation and could not easily be moved to another location, unlike the membrane system .

TABLE 2 Cost comparison of CA membranes with DEA

System size Installed capital cost Methane recovery

Membrane system

DEA

Trailor-mounted 1 .00 90%

2-Part skid 2.91 95%

Hydrotreater hydrogen recovery Oil refinery operations are requiring an increasing supply of hydrogen due to hydrotreater upgrading for the processing of heavier and more sour crude oils . Recovery of hydrogen from purge gases throughout the refinery offers a source of inexpensive hydrogen for these purposes . A typical hydrotreater schematic diagram, with integration of a hydrogen recovery unit, is shown in Fig. 4 . The hydrotreater reactor hydrogenates aromatics, olefins, sulfur and nitrogen compounds from heavy and sour crude oils . The hydrogen recovery unit processes the purge gas to concentrate the hydrogen for recycle to the make-up compressor at an interstage pressure of 250 psig (18 bar) . Pressure swing adsorption (PSA) is a suitable conventional method of recovering hydrogen for this application to compare with membrane separation . This process utilizes a bed of molecular sieves to adsorb the hydrocarbon gases to produce a 99% plus purity hydrogen gas . Upon depressurization of the sieve bed the hydrocarbons and residual H 2 are released as waste gas . A schematic diagram, showing the incorporation of PSA for H 2 recovery is shown in Fig. 5 ( 5 ] . The hydrogen recovery is 80% at a purity of 99 .5% and a pressure of 250 psig (18 bar) . The waste gas is at 5 psig (1 .4 bar) and a purity of 34% H 2 for discharge to a low-pressure burner . The schematic diagram for the membrane process is shown in Fig . 6 . In this case the hydrogen purity is less (95%) but recovery is higher (88 .5%) . As high recovery is critical and recycle purity can be as low as 93%, the membrane process therefore offers a significant process advantage over PSA .

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A relative cost comparison between the two processes for a typically sized purge gas stream is shown in Table 3 . The annual operating costs do not include membrane or sieve replacement . Over the five-year amortization time chosen it is not expected that replacement will be significant . For the PSA system all of the operating costs are associated with recompression of the product hydrogen to the hydrotreater feed pressure (56 bar) . The membrane system operating costs are higher due to the larger H 2 compression volume and to the requirement of steam for feed gas preheating . The relative hydrogen cost represents cost per unit volume of 100% hydrogen produced by each I • TO

FUEL GAS

HYDROGEN RECOVERY UNIT 800 P51G 100 °F

H2 MAKEUP

RECYCLE COMPRESSOR

MAKEUP COMPRESSOR

HYDROTREATER REACTOR

FEED

HP SEPARATOR

TO L P SEPARATOR

Fig . 4 . Integration of hydrogen recovery unit to a typical hydrotreater .

To Makeup Compressor 1st Stage 250 PSIG

34%

H2

Fig . 5 . Pressure swing adsorption unit (5 psig desorption) .

To Low Pressure Burner 5 PSIG



222

Purge Gas

72% H2

7 MMSCFD

800 PSIG

Membrane Separator (88 .5% Recovery) To Makeup 4 .69 MMSCFD 95%H2

2 .31 MMSCFD

Compressor Suction 250 PSIG

To Refinery Fuel 1. 4

25% H 2

System 60 PSIG

Fig . 6 . Separex TM membrane process for hydrotreater hydrogen recovery . TABLE 3 Cost comparison of membranes with PSA

H 2 recovery Installed capital cost Annual operating costa Hydrogen costb

Membrane system

PSA

88 .5% 1 00 1 .00 1 .00

80% 2.10 0 .81 1 .33

a Utilities based on US$ 0 05/kWh and US$ 5/MMBTU bH 2 cost based on 8000 hours per year operation for a five-year span

system . This is based on 8000 hours of operation per year for 5 years . It can be seen from Table 3 that the membrane system can recover hydrogen at a cost 25% below that of the PSA system . Comparison of two membranes for nitrogen enrichment Much research in the past has been aimed at developing membranes with high permeation rates, and this has been a worthwhile and successful endeavor One of the most permeable, if not the most permeable, membranes developed is a dimethyl silicone composite [6] . The permeation rate for a 500 A barrier layer DMS membrane is 26 .0 m3 /m2 -hr-bar compared to 0 .065 for a commercially available cellulose acetate (CA) gas separation membrane [7] . The gas selectivity for DMS is significantly lower than CA, 2 .0 for O 2 /N 2 as compared to 5 .0 for CA . However, the real effect of permeation rate and membrane selectivity can best be seen in an economic comparison for an actual application . A suitable example is nitrogen enrichment of air for use as a blanketing gas for prevention of fires around aircraft, fuel tanks and other potentially hazardous environments . Another similar case is in the production of nitrogen from turbine offgases for reinjection on offshore natural gas processing platforms to prevent subsidence of the ocean floor . A technical comparison of the DMS and CA membranes for nitrogen enrichment is shown in Figs . 7 and 8, at a nitrogen gas production rate set at



223

33,000 SCMD with a minimum purity of 95% and a minimum pressure of 1 .4 bar. It can be seen that the DMS membrane requires a significantly larger feed gas volume for compression than the CA membrane in order to produce the specified volume of enriched nitrogen . While the feed gas pressure is lower for the DMS system, the compressor horsepower is substantially greater . Additionally, it can be seen that the 0 2 purity in the permeate is considerably lower in the DMS case . All of these factors are a result of the lower selectivity for the DMS membrane . N2

Stream

33,000 SCMD 95% 7 93 bar Membrane

N2

6 .5 bar

Air Separator

308,660 SCMD 1 bar

02 Stream Compressor

275,660 SCMD 22 .8%

1428 HP

02

1 bar Fig . 7 . 0 2 /N 2 separation with DMS membrane . N2

Stream

33,000 SCMD 95% 35 5 bar

Air 62,020 SCMD 1 bar

J

Membrane

N2

32 bar

Separator 02

Stream

Compressor

29,020 SCMD 480 HP

39 .0%

02

1 bar Fig. 8 . 0 2 /N, separation with CA membrane .

A relative cost comparison between the two membranes is shown in Table 4, This will most clearly illustrate their differences as it will take into account both the higher permeation rate for DMS and the higher selectivity for CA . The capital costs are normalized to 1 .00 for the DMS membranes, including hardware and instrumentation . The operating costs are normalized to 1 .00 for the annual DMS membrane replacement cost, assuming a membrane life of 5 years for both DMS and CA . The comparison shown in Table 4 clearly demonstrates the importance of selectivity . Even with such a high permeation rate, the total DMS system

224 cost, including the compressor, is substantially higher than that of the CA system . This is a result of the need to compress a larger volume of air in order to achieve a given product volume . Operating costs are affected in the same way ; the huge compression requirements overshadow membrane replacement charges . It can also be seen that the permeate 0 2 purity is significantly less with the DMS membrane due to its lower selectivity . While this example may be a severe case of the differences seen in economics when comparing two membrane systems, it demonstrates how important it is to evaluate compression of the feed gas or permeate gas in determining which improvements in membrane performance are desirable .

TABLE 4 Membrane system cost comparison CA Membrane system cost Compressor cost

DMS

9 .20 8 .40

1 .00 22.84

Total capital cost

17 .60

23 .84

Annual operating costs Membrane repl . Power

17 00 39.25

1 .00 116 .50

Total operating costs

56.25

117.50

Conclusion It has been shown that membranes exhibit technical and economic advantages over conventional processes for many applications . Both capital and operating costs are seen to be significantly lower . When comparing one membrane system with another, however, selectivity rather than permeation rate is usually the critical factor to evaluate .

References

1 R.L . Schendel, Processing gases associated with CO, miscible flood, Gas Conditioning Conference, Norman, OK, March 7-9, 1983 . 2 W.J. Schell, C .D . Houston and W .L . Hopper, Oil Gas J ., (Aug . 15, 1983) 52 . 3 T .E . Cooley, Personal communication, 1982 . 4 Halliburton Services advertisement, Oil Gas J ., (Jan . 10, 1983) 22 . 5 R .L . Schendel, C .L. Mariz and J .Y. Mak, Hydrocarbon Process ., (August 1983) 58 . 6 R .L . Riley and R .L. Grabowsky, U .S . Patent 4,243,701, 1981 . 7 W.J. Schell and C .D . Houston, Membrane gas separations for chemical processes and energy applications, in : T.E. Whyte, C .M . Yon and E .H. Wagener (Eds.), Industrial Gas Separations, ACS Symp . Ser. No. 223, 1983, p . 125 .