Membrane Permeation Processes

Chapter 15

Membrane Permeation Processes INTRODUCTION,1238 History and Status, 1239

PROCESSTECHNOLOGY,1242 Transport Mechanisms, 1242 Design and Operating Considerations, 1245 Membrane and Module Configurations, 1246 Flow Arrangements, 1250 Simulation and Design Calculations, 1252

Fields of Application, 1258

APPLICATIONCASESTUDIES,1259 Hydrogen, 1259 Carbon Dioxide, Hydrogen Sulfide, and Water Removal, 1270 Helium Removal from Natural Gas, 1281

Air Separation, 1282 Solvent Vapors, 1288

REFERENCES,1291

INTRODUCTION Membrane technology, as applied to gases, involves the separation of individual components on the basis of the difference in their rates of permeation through a thin membrane barrier. The rate of permeation for each component is determined by the characteristics of the component, the characteristics of the membrane, and the partial pressure differential of the gaseous component across the membrane. Since separation is based on a difference in the rates of permeation rather than on an absolute barrier to one component, the recovered component that flows through the membrane (the permeate) is never 100% pure. Also, since a

1238

Membrane Permeation Processes

1239

finite partial pressure differential is required as the driving force, some portion of the permeating component remains in the residue gas, and 100% recovery is not possible. As these generalizations would suggest, the process is particularly suitable for bulk removal operations rather than for the removal of trace impurities from gas streams. It should be noted, however, that relatively high product purities and high recoveries are possible with membrane systems (at increased cost) by the use of multiple stages and recycle systems or when used in combination with other technologies. The residue gas product normally leaves the unit at a pressure close to that of the feed, while the permeate product, which must pass through the membrane, leaves at a much reduced pressure. The principal (and/or highest purity) product may be either the permeate (e.g., the production of hydrogen from dilute gas streams) or the residue gas (e.g., the purification of natural gas by the removal of excess carbon dioxide from high pressure feed), and the process may be considered either separation or purification. Gas purification and separation by membrane permeation has many advantages, including 9 Low capital investment 9 Ease of operation. Process can be operated unattended 9 Good weight and space efficiency 9 Ease of scale up. However, there is little economy of scale (see disadvantages below) 9 Minimal associated hardware 9 No moving parts 9 Ease of installation 9 Flexibility 9 Minimal utility requirements 9 Low environmental impact 9 Reliability 9 Ease of incorporation of new membrane developments. Users can install the next generation of membranes into existing equipment at the scheduled membrane replacement time (Schell, 1983) The principal disadvantages are 9 A clean feed is required. Particulates, and in most cases entrained liquids, must be removed. Filtration to remove particles down to one micron in size is preferred. 9 Because of their modular nature, there is little economy of scale associated with larger membrane installations. 9 Because membranes use pressure as the driving force of the process, there may be a considerable energy requirement for gas compression.

History and Status Since the 19th century it has been known that certain polymer membranes can separate gases by permeation. As early as 1831, Mitchell reported that different gases permeate membranes at different rates. Graham, in 1866, discussed the mechanism of permeation and demonstrated experimentally that mixtures of gases can be separated using rubber membranes. In 1950, Weller and Steiner reported on permeation processes of industrial importance, the separation of oxygen from air and the recovery of helium from natural gas. However, the selectivity and production rates of the membranes available at the time were poor

1240

Gas Purification

and the large membrane areas required made membrane permeation economically unattractive. In 1960, Loeb and Sourirajan developed a technique to cast cellulose acetate into a film that had an active thin surface layer and a highly porous supporting layer. These asymmetric membranes provided increased permeation rates while retaining their selectivity for specific gases. The development of these membranes improved the economics of membrane applications and led to increased interest in membrane technology. During the 1970s, considerable research and developmental work was devoted to membranes. Many potential applications were identified, but commercialization was slow. In 1977, Monsanto demonstrated its first full scale membrane separator at Texas City, Texas, in a hydrogen/carbon monoxide ratio adjustment application (Burmaster and Carter, 1983). In 1979, Monsanto commercialized its hollow fiber membrane module as the Prism separator. From 1979 to 1982 Prism separators were evaluated in several refinery hydrogen purification applications (Bollinger et al., 1982). The success of these pilot tests established the commercial viability of gas separation with membranes. The first large scale commercial CO2 membrane separation project was the installation of two membrane separation facilities at the Sacroc tertiary oil recovery project in West Texas in 1983. Up to 80 MMscfd of gas has been processed in these facilities (Parro, 1984). Since the early 1980s, membrane technology has advanced rapidly and continues to advance. In addition to cellulose acetate and polysulfone, the polymers used in making gas separation membranes include polyimides, polyamides, polyaramid, polydimethylsiloxane, silicon polycarbonate, neoprene, silicone rubber, and others. Today membranes can be designed to withstand a 2,000 psi pressure differential. Membranes used in hydrogen or carbon dioxide applications operate at temperatures up to 200~ while those used in solvent applications can operate at temperatures up to about 400~ (Baker, 1985). Improvements in manufacturing methods have resulted in improved membrane performance and economics. A flux increase of 5% and a separation factor increase of 20% have resulted from improved manufacturing methods (Hamaker, 1991), while during the mid1980s, air separation membranes became from two to four times more efficient. Not only are the polymers rapidly changing, but capital costs are also coming down. Advances in membranes reduced the installed cost of a membrane plant by about 40% during the 1980s (Spillman, 1989). Because of the rapid changes in technology and costs, economic studies and cost data become outdated soon after publication, and the reader should use caution in using cost data presented later in this chapter as technical advances are continuing to improve performance and reduce costs. The developmental and commercial successes of the early 1980s, and the perceived large market, attracted many companies into the field. As the technology matured and the market became extremely competitive, some companies dropped out and others changed ownership. Companies offering commercial scale membrane systems are listed in Table 15-1. The table also identifies the principal areas of application for each company's products. In the field of air separation, the improving economics of membrane-based processes have encouraged large industrial gas suppliers to join forces with membrane suppliers. This trend is pointed out by Prasad et al. (1994) who provide the following chronology: 1985: Union Carbide Industrial Gases, Inc. and Albany International Membrane Venture form a joint venture. 1986: Innovative Membrane Systems (formerly Albany International Membrane Venture) becomes a wholly owned subsidiary of Union Carbide Industrial Gases, Inc. (now Praxair, Inc.).

1241

M e m b r a n e Permeation Processes

Table 15-1 Commercial-Scale Membrane Suppliers

Application Air Company A/G Technology (AVIR) Air Products (Permea) Asahi Glass (HISEP) Cynara (Dow) Dow (Generon) DuPont Grace Membrane Systems Hoescht Celanese (Separex) International Permeation Membrane Technology and Research Nippon Kokan K.K. Osaka Gas Oxygen Enrichment Co. Perma Pure Techmashexport (USSR) Teijin Ltd. Toyobo Ube Industries Union Carbide (Linde) UOP/Union Carbide

CO2

H2

02

N2

Other*

X X

X

X X X

X X X

X

X

X

X X

X

X X

X

Note: *Includes solvent vapor recovery, dehumidification, and~or helium recovery membranes. Source: Spillman (1989) and Prasad (1994)

1988: British Oxygen Co. (BOC) and Dow/Generon Membrane Systems form a joint venture. 1989: L'Air Liquide and DuPont form a joint venture. 1990: Medal AL becomes a wholly owned subsidiary of L' Air Liquide, Inc. 1991: Permea, Inc. becomes a wholly owned subsidiary of Air Products and Chemicals, Inc. The commercial application of membrane-based hydrogen processes also expanded rapidly because of the high permeability of selected membranes for hydrogen, the high value of pure hydrogen, and the wide range of hydrogen containing gas streams in refinery, petrochemical, and industrial chemical operations. According to Koros and Fleming (1993), the major suppliers of membrane systems for hydrogen applications are Permea, Medal, UOP, Ube, and Separex. In 1996, Medal reported that more than 60 of their hydrogen recovery systems were in operation or under construction (Medal, 1994).

1242

Gas Purification

PROCESSTECHNOLOGY Transport Mechanisms It is generally agreed that a solution-diffusion mechanism governs the transport of gases through all commercially important nonporous membranes. The mechanism involves the following: (a) adsorption of the gas at one surface of the membrane, (b) solution of the gas into the membrane, (c) diffusion of the gas through the membrane, (d) release of the gas from solution at the opposite surface, and (e) desorption of the gas from the surface. Since these steps are not necessarily independent, the term permeation is used to describe the overall transport of gases through a membrane. In its simplest form, the solution-diffusion model considers only steps b, c, and d. This model is based on two assumptions: (1) the concentration of a component in a membrane at its surface is directly proportional to the partial pressure of the component in the gas phase adjacent to the surface, and (2) the rate at which a component passes through a membrane is proportional to the concentration gradient (concentration/distance) in the membrane. These two assumptions represent Henry's law and Fick's first law of diffusion, respectively, and can be stated as follows: c i - kip i

(Henry' s law)

(15-1)

Ji -"-Di(dci/dx)

(Fick's law)

(15-2)

Where: c i -- local concentration of i in the membrane Ji - steady state flux of i k i = solubility coefficient Pi = partial pressure of i in the gas D i = local diffusivity x = distance through active membrane Combining and integrating equations 15-1 and 15-2 over the full membrane thickness across the membrane, yields Ji = PiApi//

(15-3)

Similarly, for a hollow tubular membrane, such as a hollow fiber, the steady state rate of gas permeation is (Stem, 1986) Ji = Pi(2rtLApi/ln(Ro/Ri)) Where:

Pi = Api = 1= Ro = Ri = L=

(15-4)

kiDi, the permeability coefficient Pi(feed)- Pi(permeate) (the partial pressure difference across the membrane) membrane thickness (x = 1) effective outer radius of the tube effective inner radius of the tube length of tube

Membrane Permeation Processes

1243

Any consistent set of units may be used in equations 15-1-15-4. Permeability coefficient data are often given in Barrers. One Barrer = 10-1~ 3{STP })(cm)/(cm2)(sec)(mmHg). Since commercial membranes normally consist of a very thin active layer on a thicker porous substrate, the effective thickness may not be accurately known, and it is more convenient to use the permeation rate, Pi/l, for the overall membrane, as a correlating factor. Typical engineering units for the permeation rate are (scf)/(ftZ)(hr)(100 psi). One (scf)/(ftZ)(hr)(100 p s i ) - 1.55 • 10-5 (cm3){ STP})/(cm2)(sec)(cmHg). The units of flux, J, are (scf)/(ftZ)(hr). The Henry's law/Fick's law model previously described is a simplification of the actual permeation mechanism, and more complex models have been proposed. The dual system model, for example, is a more precise representation for many cases. It assumes that gas molecules which dissolve in the dense regions of the membrane surface follow Henry's law; while molecules that adsorb on the walls of microscopic cavities in the membrane surface follow Langmuir's adsorption isotherms. Equations based on the dual system model have been developed and presented by Lee et al. (1988). Additional discussions of transport mechanisms are provided by Lacey and Loeb (1972) and Stern (1986). Permeation rates for specific components and membranes are not constants. They vary with temperature, pressure, and the presence of other components in the gas. Detailed permeation rate data for commercial membranes are normally considered proprietary; however, some comparative data have been published. Table 15-2 lists typical permeation rate data for a number of membranes and gases. The cellulose acetate data are from Mazur and Chan (1982), and the other data are based on a paper by Tomlinson and Finn (1990). The permeation data given in Table 15-2 are relative values based on a permeation rate of 1.0 for oxygen in cellulose acetate and polysulfone. However, data presented by Schell and Hoernschemeyer (1982) for the permeation rates of various gases in cellulose acetate indicate that actual values, expressed as (scf)/(ft2)(hr)(100 psi), are close to the relative values listed in the table. For example, they show the permeation rate for oxygen, measured at room temperature and 100 psig, to actually be about 1.0 (scf)/(ftZ)(hr)(100 psi), and the actual permeation rates for other gases in cellulose acetate to be similar to those listed.

Table 15-2 Permeation Rates of Gases Through Membranes Relative Permeation Rate (1) Membrane

Hz

Polysulfone Cellulose Acetate Polyamide Dow Product Permea Product PDMS (2)

13 12 9 136 22 649

N2 0.2 0.18 0.05 8 0.4 281

02

CH4

COz

H20

He

HzS

CO

C2H6

1 1 0.5 32 2.3 604

0.22 0.2 0.05

6 6

100

15

10

0.3

0.1

0.4

93 9

Notes: 1. Permeation rates are relative, based on 1 for oxygen in polysulfone and cellulose acetate. 2. PDMS = poly(dimethyl siloxane); silicone rubber. Sources: Tomlinson and Finn (1990) and Mazur and Chan (1982)

1244

Gas Purification

Lee et al. (1995) evaluated available data on the permeability of gases in cellulose acetate to obtain values for use in designing systems for removing CO2 from natural gas. They concluded that the following would be reasonable values for computer simulation of the process: Permeability rate for C O 2 : 9 • 10-5 (cm3){ STP})/(cm2)(sec)(cmHg) (or approximately 5.8 (scf)/(ftZ)(hr)(100 psi) Selectivity for CO2/CH 4 = 20 Selectivity for Nz/CH 4 = 1 Selectivity for (C2+)/CH 4 -- 0.4

Additional permeability data for cellulose acetate membranes are provided by Li et al. (1990), Donohue et al. (1989), and Ettouney et al. (1995). Gases with high permeability rates are often called "fast" gases. The data in Table 15-2 indicate that low molecular weight and highly polar gases tend to be fast gases; while the slower gases are nonpolar and/or higher molecular weight. Detailed data on organic solvent permeation rates are given by Baker et al. (1986). The separation factor of a membrane, c~ij, is defined as: Oqj : Pi/Pj

(15-5)

and is an indication of a membrane's ability to separate species i and j. A separation factor greater than 1 indicates that, at equal partial pressures, component i permeates through the membrane faster than j. Very high (or very low) separation factors result in easy separations. No separation is possible if oqj = 1. Separation factors are influenced by membrane materials, feed composition, temperature, and pressure. Separation factors for various systems are presented in Table 15-3. Note that the separation factor for oxygen and nitrogen is considerably lower than those for other component pairs generally considered viable for commercial membrane separation. Oxygen and nitrogen are very similar in molecule size and solubility, which makes their separation via

Table 15-3 Separation Factors for Various ComponentSystems in Commercial Membranes

Component System CO2/CH 4

O2/N 2 H2/CH4 H2/CO Hz]N 2 HzS/CH 4

He/CH4 CO2/C2H6

Ranges of Separation Factors

Typical Separation Factors in Cellulose Acetate

10-50 3-12 45-200 35-80 45-200 40-60 60-100 44-52

25

Sources: Stookey et al. (1986) and Spillman et al. (1988)

45 45 50 60 50

Membrane Permeation Processes

1245

membranes one of the more difficult processes. However, due to the abundance of "free" feedstock a high recovery is not required. The separation factors for solvent/air processes have a very wide range depending on the membrane material. For example, the separation factor for toluene/nitrogen can be as low as 40 for nitrile rubber or as high as 10,000 for neoprene (Baker et al., 1986).

Design and Operating Considerations The separation efficiency of a membrane for a given gas mixture will depend on the gas composition, the pressure difference between the feed and the permeate, and the separation factor for the two components at the specific operating conditions. The higher the separation factor, the greater the selectivity of the membrane and the higher the product purity. The gas composition and pressure differential become very important when the more permeable gas in the feed has a low concentration. Since the partial pressure of the fast component on the permeate side cannot exceed its partial pressure on the feed side, high feed gas pressures and low permeate pressures are required to obtain efficient separations even with high separation factors. The differential pressure across the membrane relates directly to the membrane area required. Compression costs on the other hand are a function of pressure ratio. Therefore, operation at high pressure with a substantial pressure differential across the membrane but with a reasonably low pressure ratio, is economically advantageous where recompression of the permeate is required. The general effects of varying key operating factors, while holding other conditions constant, for a typical single-stage membrane permeation system, can be summarized as follows: 1) Increasing the overall differential pressure across the membrane leads to an increase in permeate flow rate and a decrease in the concentration of the fast gas in the permeate stream. The differential pressure at any point in a module can be affected by pressure drop in either the feed or permeate flow channels. If the permeate flow rate is sufficiently high it can generate a back pressure that reduces the differential pressure and rate of permeation. 2) Increasing the feed gas flow rate decreases the percent recovery of the fast gas as permeate and decreases the purity of the residue. However, increasing the feed rate increases the purity of the permeate and increases the percent recovery of the slow gas in the residue. 3) Decreasing the feed gas flow rate below a critical value decreases the separation efficiency due to a boundary layer effect. The concentration of the fast gas is depleted in the feed gas adjacent to the membrane surface, reducing its partial pressure and therefore its rate of permeation. Since the rate of permeation of the slow gas is not affected (or is increased) this reduces permeate purity. The critical flow rate is determined by the degree of mixing at the membrane surface, and this is a function of gas velocity; gas properties such as viscosity, density, and diffusivity; and module design. 4) Increasing the temperature raises most permeabilities by about 10 to 15% per 10~ and has little effect on separation factors (Schell and Hoernschemeyer, 1982). 5) Increasing the membrane area increases the purity of the residue; while decreasing the membrane area increases the purity of the permeate. In hydrocarbon membrane systems, permeation of fast gases (e.g., H 2 or CO2) increases the concentration of heavy components (e.g., C3+ hydrocarbons) in the remaining gas. The increased concentration of readily condensible components, as well as the decrease in temperature that frequently accompanies permeation, may cause the condensation of liquid

1246

Gas Purification

hydrocarbons within the membrane unit, interfering with its operation. Typical remedies for this problem are preheating the gas and removing heavy hydrocarbons from the gas ahead of the membrane unit. Even where membranes can physically tolerate the condensate, performance will normally suffer. The condensed liquid (usually made up primarily of slow permeating components) can cover the membrane surface, forming an additional barrier to permeation. Alternatively, the liquid may wet the membrane creating a gas leakage path, which can result in permeate contamination. Membrane life is an important factor affecting process economics. Membranes typically require replacement every three to seven years. It is therefore necessary to consider factors that may shorten or possibly extend their life. Funk et al. (1986) investigated the effect of impurities in the gas on a cellulose acetate membrane in acid gas service. They found that various components in the gas affect the permeability, the tensile strength, and the elastic modulus of the membrane. Some of the adverse effects of individual components can be mitigated by proper design (e.g., temperature control), but other impurities may require removal prior to membrane processing. Therefore, the feed gas should be evaluated for the presence of particulates, entrained liquids, oil mist (compressors), condensible components, all trace compounds, and water. Pretreatment systems that may be needed include: l) A high efficiency separator to remove particles and oil mist 2) A liquid knock-out drum to remove liquid hydrocarbons and water 3) Preheat or reheat steps to raise the gas temperature sufficiently above its water and hydrocarbon dew point to prevent condensation in the module 4) Component removal steps to eliminate compounds, such as solvents, BTX, or ammonia, that may cause membrane deterioration The design of pretreatment equipment must also account for both normal operating and upset conditions. Since many of the process variables are highly interdependent, the overall design of membrane systems requires consideration of all requirements and conditions. This is illustrated in Figure 15-1, which shows the effects of percent hydrogen in the permeate, percent hydrogen recovery, and permeate pressure on the relative costs of recovering hydrogen in a typical application. These three factors, as well as others, must be evaluated to optimize overall economics for each specific case. Most membrane system suppliers have computer programs to optimize the system design (Poffenbarger and Gastinne, 1989; Spillman, 1989; MacLean et al., 1983).

Membrane and Module Configurations Membranes The key requirements for a membrane to be used in an economical gas purification or separation process are: 1) High permeability for the component to be removed 2) High selectivity for the component to be removed in relation to other components 3) High membrane stability in the presence of all gas components which will come into contact with the membrane 4) Uniformitywfreedom from pinholes or other defects

Membrane Permeation Processes

1247

1.25 -

% Hz Recove r y 0

1.15 95

> 0 o

40

~'~.,,..,~0

-

%%

7,,,

%

ff

%

"'-C'-._ -,_ -,. " - , _ ' . . / I "-',,,~>.. -"'.,. -"',,, " , 1 "i

o

~ 1.05 m @ rr

1 4 0 psia 0.95

I 90

I

I I I 1 I 95 % H y d r o g e n in P e r m e a t e

J

!

I 1 O0

Figure 15-1. Effects of hydrogen permeate purity, permeate pressure, and percent hydrogen recovery on relative cost of hydrogen recovery from a 75% H2, 815 psia refinery gas stream. (/-leyd, 1 9 8 6 )

5) Low effective thickness of the active portion of the membrane to ensure a high permeation rate 6) Physical strength to withstand the required operating conditions Items 1 through 3 are related primarily to the polymers used in fabricating the membrane; while items 4 through 6 relate to the fabrication method. A key breakthrough in the development of fabrication methods was provided by Loeb and Sourirajan (1960). They developed a technique for casting asymmetric cellulose acetate membranes with a uniform, very thin (0.1-1.0 micron) skin on a strong porous substrate (100-200 microns thick). Since this original work, which was actually aimed at the development of reverse osmosis membranes, the basic approach has been applied to a variety of polymeric materials and to sheet and hollow fiber configurations for use in both gas and liquid phase operations. Development work has also investigated alternative asymmetric membrane systems, including (a) an ultra thin nonporous film laminated to a much thicker microporous backing (which may be a different material) and (b) a very thin nonporous film applied as a coating to a thicker microporous substrate (Stern, 1986). A complex membrane structure reportedly used in the Monsanto Prism separator is a "skinned" asymmetric hollow fiber of polysulfone coated with a thin film of silicone rubber (about 1 micron thick). The polysulfone skin (about 0.1 micron thick) is the active separator, while the silicone rubber serves to seal any defects in the base membrane without affecting the intrinsic permeability of the membrane (Koros and Chern, 1987). Koros and Fleming (1993) present a comprehensive review of membrane-based gas separation with emphasis on membrane materials, formation techniques, and module designs. The most popular module configurations are the hollow fiber and spiral-wound designs due to their high packing density.

1248

Gas Purification

Spiral-Wound Configuration In the spiral-wound configuration an envelope is formed with two membrane sheets separated by a porous support material. Typically the module consists of several such envelopes. The material between the membranes (permeate channel spacer) supports them against the operating pressure and defines the permeate flow channel. The envelope is sealed on three sides. The fourth side is sealed to a perforated permeate collection tube, and the envelope is wrapped around the collection tube with a net-like spacer sheet that has two functions: 1) It keeps adjacent membranes apart to form a feed channel. 2) It promotes turbulence of the feed gas mixture as it passes through the module, thus reducing concentration polarization. During operation, the feed gas mixture enters one face of the module, travels axially along the feed channel spacer and membrane surface, and exits the other face as a residue or retentate. The more permeable gases pass through the membranes and travel in a spiral path inward within the envelope through the permeate channel spacer until they reach the perforated collection tube and finally exit as permeate (Figure 15-2). The feed channel spacer is a key feature of the spiral-wound module and, as shown by Da Costa et al. (1991, 1994), its design significantly affects module performance. Typically the modules have about 1,000 square feet of surface per cubic foot of volume and are 4-12 inches in diameter by 36-42 inches long. Up to six modules may be housed in a single pressure vessel shell.

Hollow Fiber Configuration The hollow fiber configuration consists of thousands of hollow fibers packaged in bundles mounted in a pressure vessel resembling a shell and tube heat exchanger. For high pressure applications the fiber diameter is usually on the order of 100 ~tm ID and 150-200 ~tm OD. The bundles are capped on one end and have an open tube sheet on the other end

Figure 15-2. Diagram of a spiral-wound membrane permeation element. (Courtesyof

Membrane Technologyand Research, Inc.)

Membrane Permeation Processes

1249

(Figure 15-3). The bundles typically have 3,000 square feet of membrane surface per cubic foot of module volume and the module dimensions range from 4-12 in. in diameter by 4-20 ft long. The feed gas is introduced on the shell side because hollow fibers are much stronger under compression than expansion. The faster permeating gases migrate into the fiber bore and exit via the open end of the bundle. For low pressure applications the fibers have a diameter greater than 400 lam and the feed gas enters the bore side while the permeate exits via the shell side. This configuration reduces pressure drop on the feed side. Not all membrane materials can be made into a thin selective layer on a porous substrate in a hollow fiber form. Consequently, spiral-wound membranes, which can be made from a wider range of materials, usually have higher permeation rates. However, this is offset by the much higher packing density of hollow fiber modules, resulting in similar overall productivity per unit module volume for the two configurations. This situation could change if developments in polymer science lead to more effective thin films in a hollow fiber form. Numerous mechanical designs of modules have been developed for both the spiral-wound and hollow-fiber concepts. The various designs are aimed at optimizing such features as: membrane area per unit volume, gas flow distribution and pressure drop, seals and fastenings, and assembly technology. With either spiral-wound or hollow-fiber systems, large

Figure 15-3. Diagram of a hollow-fiber membrane permeation element. (Courtesyof

Permea)

1250

Gas Purification

commercial installations normally require a large number of individual modules. This is evident in Figure 15-4, which shows a UOP Advanced Membrane System for removing carbon dioxide from natural gas. A third module design--plate and framehutilizes a stack of disk shaped membranes separated by sheets of porous filter paper and membrane support plates. The flow pattern is very much like that of a plate and frame filter. This design is not competitive for large commercial applications because of the relatively small membrane area per unit volume attainable. However, plate and frame designs are used for producing oxygen enriched air in small medical applications.

FlowArrangements To obtain the specified flowrates and purities, the optimum arrangement of the modules needs to be determined. Series flow, Figure 15-5, provides for high recoveries for a given feed rate. The permeate purity varies from module to module, which makes it possible to produce multiple permeate products. The substantial velocity changes that occur when large portions of the feed are recovered as permeate can be accommodated by using successively smaller elements. The series flow arrangement has a slightly higher pressure drop between the feed and the residue than a parallel flow arrangement. Parallel flow, Figure 15-6, allows for higher feed rates for the same recovery. Turn down is accomplished by taking elements out of service. In designing parallel flow systems, particular care must be paid to the gas distribution systems. Non-uniform gas flow to the elements

Figure 15-4. Large commercial installation of UOP Advanced Membrane System for removing carbon dioxide from natural gas. (Courtesyof UOt~

Membrane Permeation Processes

1251

Residue Gas (Non-Permeate Gas)

Feed Gas

Permeate Gas Figure 15-5. Series flow configuration.

Residue Gas (Non-Permeate Gas)

Feed Gas

Permeate Gas Figure 15-6. Parallel flow configuration.

can result in declines in product purity and recovery. High recovery in one element can cause liquid formation. The combined series-parallel flow sequence, Figure 15-7, provides for good turndown ratios while accommodating high feed rates with low residue flow rates. A reduction in the number of modules in parallel accommodates the reduction in flow as permeate is removed. As stated earlier, separation by membrane permeation is not an absolute separation. Each species has a finite permeability through the membrane and the enrichment is achieved due to relative permeabilities, not zero permeability for one of the species. As higher purities are approached, recovery of product declines rapidly and single-stage systems become increasingly inefficient. Therefore, single-stage processes frequently are applicable to bulk removal of a species or the concentration of feeds to other purification processes.

1252

Gas Purification

Feed Gas

Residue Gas (Non-Permeate Gas)

Permeate Gas

Figure 15-7. Combined series-parallel flow configuration. As an alternative, multi-stage membrane systems with recompression arrangements may be employed. Multi-stage arrangements, Figure 15-8, work well when high permeate purity or improved recovery rates are required. Additional multi-stage arrangements are described by Spillman et al. (1988).

Simulation and Design Calculations In considering membranes for various applications, the engineer must have a means of estimating the performance of a membrane system. The fundamental laws of diffusion discussed earlier in this chapter under "Transport Mechanisms" apply. Mathematical solutions for compositions of residue and permeate streams have been developed for two components and are described in the literature (e.g., Lacey and Loeb, 1979). For three or more compo-

9= mr

Residue

Feed

I Compressor

T

Figure 15-8. Multi-stage flow configuration.

Permeate

Membrane Permeation Processes

1253

nents a strict mathematical solution to the differential equations has not been found. However, a numerical solution is possible, and with the prevalence of computers, this approach is not only practical but widely used. The basic equation defining the rate of flow of a component through a membrane is equation 15-3 which can be written in the form: Ji = Ri (Pi,feed- Pi,perm) Where:

(15-6)

Ji = steady state flux of i, moles/(time)(area) = Mi/A M i = flow rate of i through a given area, moles/time kiDi Pi Ri . . . . l 1 Pi = partial pressure of i in the gas A = area k i - solubility coefficient for i D i = local diffusivity for i l = membrane thickness Pi - permeability coefficient for i

Any consistent set of units may be used. Gas volume (e.g., scf) is often used instead of moles. The partial pressure of each individual component changes as the separation is carried out because different components are being removed from the feed (high pressure) side at different rates. Because the partial pressure of individual components is a function of position along the membrane surface, mathematical integration of equation 15-6 over the entire length of a membrane surface is an interesting exercise for two components, but not practical for three or more components. Commercial membranes today are asymmetric membranes. This means that the active membrane surface is only a very thin layer on top of a porous substrate. Components which permeate the membrane must travel out of the porous membrane substrate before entering the bulk permeate stream. Therefore, the effective partial pressure and mole fraction on the permeate side of the active membrane surface are only a function of the material passing through the membrane, not a function of the bulk permeate stream. A computer simulation of the process can be made by considering small incremental areas of the membrane individually. A permeation analysis and a material balance are performed on the first incremental area. The residue gas from this area is treated as the feed to the next area and the operation is repeated. The analysis continues, adding areas until the residue gas meets the product purity requirement or other process requirements are attained. If the assumed areas are small enough, the feed and residue gas compositions for each AA are similar enough that the driving force for permeation can be based on the composition of the feed rather than on an average of the feed and residue compositions. The permeate composition for each incremental area must take into account the permeation rates of all components, but is not affected by permeate from the other areas. Figure 15-9 is a simple diagram showing gas flows adjacent to and through an increment of area, AA. The computer simulation is based on this diagram and the following equations: For a small increment of membrane area, equation 15-6 becomes Mi = Ri (Pi,feed -- Pi,perm)AA

(15-7)

1254

Gas Purification ~-- AA, Small Increm( =rement of Area Residue Gas

cFeed Gas '

NI

%l

N'

v

rZ///////,;,

.

Active Membrane Porous Substrate

.. k J . Iv

01

W '~

O=

"= Bulk Permeate Ir

N, = Flow of Feed to AA N= = Flow of Residue from AA = Flow of Feed to Next AA O1 = Bulk Flow of Permeate Prior to AA O2 = Bulk Flow of Permeate After AA M = Flow of Permeate Through AA

Figure 15-9. Diagram of gas flow pattern at small increment of membrane area for computer simulation model.

Substituting for partial pressures: Mi = Ri (Yi,feedgreed- Yi, perm 7tperm)AA

(15-8)

The mole fraction of any component, i, in the permeate is Yi,perm= Mi/]~Mi

(15-9)

Similarly, the mole fraction of component i in the residue from increment area AA is Yi,residue= N2i/]~N2i

(15-10)

Combining equations 15-8 and 15-9 yields Mi

Ri AA y i,feed n feed

(15-11)

I 1+ RiAAzMi~permI

This is the key equation in the computer simulation program. Other important equations define the increase in bulk permeate and decrease in residue gas flow at each AA: 02, i --01, i + M i

(15-12)

N2,i = N1,i - M i

(15-13)

Membrane Permeation Processes

1255

In equations 15-7 through 15-13, AA = incremental area M i - flow of component i through AA, moles/time • M i = f l o w o f all components through AA N1 = flow of feed gas to AA N 2 = flow of residue gas from AA Ol = bulk flow of permeate from prior AA' s O2 = bulk flow of permeate after AA Ri = permeability rate for i, moles/(time)(area)(partial pressure differential) Yi = mole fraction i in gas 7r = total pressure The permeation rate for each component i must be known, and the total feed pressure and permeate pressure must also be known. A reasonable ZM i is assumed for the first increment of area, AA. It is then corrected by subsequent iterations, until a solution is found where M i calculated for each component equals (within the specified tolerance) the M i used in the previous iteration. The sum of all M i from each iteration is also used for the calculation of the new ~M i. At the first increment of membrane area, Yi,feedis simply the mole fraction of component i in the feed. After the solution for M i for each component has converged, the residue gas molar flow, N2,i, for each component is calculated by subtracting Mi, the permeate flow, from the feed flow N l,i. This residue flow then becomes the feed flow to the next small increment of membrane area. The procedure is repeated until the desired product specifications are met. After each increment of area is calculated, the permeate flow for each component, Ol,i, is increased by the permeate flow for that incremental area, M i. Similarly, at each increment of area, the residue gas for each component, N l, is reduced by the permeate flow passing through that increment of area, M i. The sum of all N2,i's from the final increment of membrane area is the residue gas for the system. Each increment of area is also accumulated so that the total area required is known as well as the gas compositions and quantities of the residue and permeate streams. A simplified flow sheet for the required calculations is shown in Figure 15-10. Table 15-4 summarizes computer output from a program using the calculation logic outlined in Figure 15-10. The data are from a computer simulation of a spiral-wound cellulose acetate membrane unit recovering CO2 from 7.4 MMscfd (812 mol/hr) of a high COz content gas, rich in hydrocarbons, as might be found in a CO2 flood enhanced oil recovery (EOR) project. Input to the program consisted of the feed composition, flow rate, pressure, and temperature; the permeate pressure; and the specification that the residue gas contain 40 vol% C O 2.

In addition to the flow rates and compositions of the residue gas and permeate, the output indicates the approximate membrane surface area required, and the number of increments of membrane area used for the calculation. At 1,000 sq ft of membrane surface area/cubic foot of membrane module volume, approximately 13.1 c u f t of membrane module volume is required. A membrane module 8 in. in diameter and 36 in. long occupies approximately 1 cu ft and contains approximately 1,000 sq ft of membrane surface area. Therefore, the simulation indicates that 13 modules (8 in. diam x 36 in.) are required for the separation.

1256

Gas Purification

INPUTSYSTEMDATAAND PRODUCTSPECS

INITIALJZEFORSMALLINCREMENTOFAREA,&A. SET IK FOREACHCOMPONENTATA REASONABLEVALUE e.g., Mi - Rte~La.eYi,Feed(~feed" ~perm /2)

CALCULATEINITIALT..,Mi FROMINITIALM FOREACHCOMPONENT

CALCULATENEW ~ FFOREACHCOMPONENTBYEQUATION15-11

NEW M,- OLDM,? ~ _

i.e., is ...NEW.IK- OLDI~ < 0.0001 ? .._ NEW M, _~

ADD NEW M~TO PRODUCENEW~-,M REPEATFOREACHCOMPONENT

IS FLAGSET FORANYCOMPONENT?

NO

REPEAl"FOREACHCOMPONENT

INCREASEFLOWOFPERMEATEBYFLOWTHROUGH&A, USINGEQUATION15-12

REDUCEFLOWOFPERMEATEBYFLOWTHROUGH&A, USINGEQUATION15-13 REPEAl"FOREACHCOMPONENT

CALCULATENEWY. RESIDUEFROM (EQUALSY,= TO NEXT~uz.), EQUATION15-10 REPEATFOREACHCOMPONENT

NO

m m = ~ _

~

-

-

ARESPECSMET?

J~rou4 INCREASE~ BYNEWINCREMENTOFAREA

RESULTSJ

(ENO) I

USENEWY,FEED(YiRESIDUEFROMPREVIOUSAREA) AND INITIALIZEWITH fiNALT..M,FROMPREVIOUSITERATION

Figure 15-10. Calculation flowsheet for computer simulation of membrane permeation.

1257

M e m b r a n e Permeation Processes

Table 15-4 Computer Simulation of C02 Recovery Using a Single-Stage, Spiral-Wound Cellulose Acetate Membrane Unit

Component

MOLl HR

Feed MOL FRAC.

Residue MOLl MOL HR FRAC.

Permeate MOLl MOL HR FRAC.

Carbon dioxide Nitrogen Methane Ethane Hydrogen sulfide Propane Isobutane Butane Isopentane Pentane Hexane Heptane

735.58 1.32 40.24 11.74 .10 12.41 2.28 4.71 1.41 1.34 .46 .41

.905887 .001626 .049557 .014458 .000123 .015283 .002808 .005800 .001736 .001650 .000567 .000505

44.29 1.12 33.17 10.65 .00 11.63 2.14 4.41 1.32 1.26 .43 .38

691.29 .20 7.07 1.09 .10 .78 .14 .30 .09 .08 .03 .03

TOTAL

812.00

.399763 .010074 .299400 .096103 .000006 .104938 .019279 .039827 .011923 .011331 .003890 .003467

110.78

.985849 .000291 .010084 .001559 .000142 .001119 .000206 .000425 .000127 .000121 .000041 .000037

701.22

Area = 13,110.9 Square feet, 125 Increments used Note: Operating Conditions: Feed at 500.0 psia, 140.0 degrees F, permeate 100.0 psia. Source: Schendel (1995)

The permeation rate, R i, is not necessarily a constant. For example, it is known to vary with temperature and pressure, and corrections for these factors are readily included in the calculations. Because changes in composition from increment to increment are normally small, even component interference effects can be included in the calculations if required. Gas composition effects can be significant. The concentration of carbon dioxide in natural gas, for example, can affect its rate of permeation (per unit of partial pressure differential). Data provided by Hogsett and Mazur (1983) indicate that increasing the CO2 concentration from 15 to 50% in 200 psia natural gas increases the CO2 permeation rate from about 4.0 to 5.0 (scf)/(ftZ)(hr)(100 psi) with GASEP membranes. Hogsett and Mazur (1983) also suggest a simplified approach for estimating the approximate membrane area required for a multicomponent system. The approach avoids the need to use a computer simulation model, but is reportedly accurate to only about _+20% of the actual required area. The method is based on the following steps: 1) Sort all components into two groups (fast and slow permeators), with the fast permeators typically having permeation rates 15 times those of the slow permeators. 2) Calculate weighted average permeation rates for the two groups. 3) Use equations developed for two-component systems to calculate the membrane area required to meet fast permeator removal requirements.

1258

Gas Purification

The required two-component equations and details of the calculation procedure are given by Hogsett and Mazur (1983). Ettouney et al. (1995) compared two-component and four-component models for the analysis of natural gas well enrichment. In the flow system evaluated, gas from the well is recycled through a permeation unit and back to the well. Removal of a CO2 and HzS-rich permeate results in an increase in purity of the well contents with time. Results of the study showed that proper design of this system requires that the models take into consideration flow patterns, the presence of more than one species, and permeability rate functions that include the effects of both composition and pressure. The use of a simplified two-component model with constant permeabilities gave large deviations from the more detailed models.

Fields of Application Membrane systems are now available that are economically attractive for many applications, and the fields of application are growing steadily. The main current commercial applications of membrane-based gas purification and separation are: 1) Hydrogen recovery from nitrogen-bearing gases, e.g., ammonia synthesis purge gas 2) Hydrogen removal and recovery from hydrocarbons (e.g., methane) and other slower permeating gases (e.g., carbon monoxide) 3) Removal of carbon dioxide, hydrogen sulfide, and water vapor from methane and other hydrocarbon gases (e.g., upgrading natural gas to meet pipeline specifications) 4) Oxygen or nitrogen separation from air (e.g., producing nitrogen for inert blanketing) 5) Helium removal and recovery from natural gas 6) Solvent vapor removal from exhaust gases The small molecular size of hydrogen and helium allow them to diffuse rapidly through membranes. Therefore, these elements are readily separable from larger molecule gases such as methane and heavier hydrocarbons. The acid gases, carbon dioxide and hydrogen sulfide, and water have larger molecules and diffuse more slowly than hydrogen or helium. However, they are much more soluble in polymers used for the manufacture of membranes. Since permeability is the product of solubility and diffusivity, highly soluble gases can have permeation rates comparable to those of lighter gases; while gases such as methane, which has both a relatively large molecule and low solubility, have low permeation rates. The separation of oxygen and nitrogen is difficult because the size and shape (and hence the diffusivity) of the molecules are quite similar. In addition, the solubility and diffusivity of the faster gas (oxygen) are generally quite low, resulting in a low rate of permeation and the requirement for a large membrane area. However, because of the industrial importance of this separation and the scarcity of simple alternatives, it has been the subject of extensive research and development work. In addition, the feedstock is "free." Therefore, high recovery is not a requirement for this separation. Low recovery operation can be used to improve the separation efficiency (i.e., the permeate purity). Organic solvents typically exhibit low diffusivity rates but very high solubilities in appropriate membranes. As a result, satisfactory permeation rates can be obtained relative to the other components in exhaust gases. The partial pressure differential can be improved by operating the permeate side under vacuum. This is normally more economical than compressing large volumes of atmospheric pressure exhaust air. However, the process is not effective in the extremely low partial pressure range required for removing traces of organic

Membrane Permeation Processes

1259

solvents from air, and is used primarily for the bulk removal of solvents from relatively concentrated exhaust air streams.

APPLICATIONCASESTUDIES Hydrogen Ammonia SynthesisPurge Ammonia is produced by reacting hydrogen with nitrogen over a catalyst. The hydrogen is usually produced in a steam-methane reformer (SMR) and the nitrogen comes from the air supplied to a secondary reformer. Not all the methane (natural gas) is converted to hydrogen in the SMR nor is the conversion of nitrogen and hydrogen to ammonia complete in the synthesis reactor. The reactor product gases, after removal of the ammonia, are recycled back to the reactor feed to improve yields. Inert gases, such as argon from the air and methane, build up in this closed loop and reduce the nitrogen-hydrogen reaction rate. Therefore, a continuous gas purge to the fuel gas system is maintained to keep these inerts at a manageable level. However, this purge gas contains valuable hydrogen. Currently many ammonia plants recover hydrogen from the purge stream. This is a popular application for membranes because of the high permeation rate of hydrogen and the high pressure of the purge gas. Shirley and Borzik (1982) reported on the installation of a hydrogen recovery membrane system in a 1,000 ton per day ammonia plant that resulted in a 5% overall capacity increase. MacLean et al. (1988) reported the installation of a membrane unit in a 600 ton per day ammonia plant. The unit provided an 89% hydrogen stream with an 86% hydrogen recovery, representing a 4% ammonia plant capacity increase. A comparison of membrane separation versus cryogenic separation in a typical large ammonia plant was made by Schendel et al. (1983). For the case study described in this paper, the operating conditions were modified to increase the methane content at the entrance to the synthesis loop to about double that allowed without hydrogen recovery. The pertinent process variables for the 15 MMscfd feed stream are summarized in Table 15-5. To prevent densification of the membrane or formation of an insoluble phase in the cryogenic system, the ammonia in the feed to both systems is reduced to very low levels in a water scrubber. To prevent formation of a solid phase in the cryogenic unit, molecular sieves are used to remove the water picked up in the scrubber. The inerts concentration in the synthesis loop is held constant. Since the cryogenic system produces a slightly purer recycle hydrogen stream, the purge rate required to hold the inerts at a fixed level is less in this system. Hydrogen permeates through the membrane and is recovered as a low pressure product. The recompression requirements of the membrane system are, therefore, considerably greater than those of the cryogenic unit. To reduce the recompression costs, the membrane unit is operated in two stages. The utility requirements are also presented in Table 15-5. The external power requirement for the compressor horsepower is that required to integrate the recovery unit into the total plant. The primary advantage of the cryogenic unit is the lower recycle recompression requirements. The higher on-skid electrical costs for the cryogenic unit reflects the use of the molecular sieve to dry the feed stream to the recovery unit. The cryogenic unit also requires a small purge stream of nitrogen. The estimated capital equipment costs are virtually identical, approximately $1.35 million (1983 dollars). The membrane unit consists of two skids, while the cryogenic unit requires

1260

Gas Purification

Table 15-5 Comparison of Membrane and Cryogenic Separation Units for Hydrogen Recovery in an Ammonia Plant

Membrane Type, Monsanto PRISM Feed Composition, mole% H2 N2 CH 4 Ar NH3 Feed Gas Quantity, lb moles/hr Pressure at Separator Inlet, psig Hydrogen Recovery, % Hydrogen Purity, mole% Ammonia Recovery, % Recycle Product to High Stage Compressor, % Recycle Product to Low Stage Compressor, % Electricity, kWh~ Cooling water, gpm Steam (600 psig), lbs/hr Nitrogen, scfh Instrument air, scfh Turbine condensate, gpm External power differential, kWh/h

60.8 20.0 12.1 3.2 3.9 1,767 1,973 95.7 87.8 99.8 49.8 50.2 30 200 1,910 Startup only 2,100 12 470

Cryogenic Type, Petrocarbon S-2000 60.8 20.0 12.2 3.1 3.9 1,503 1,000 94.6 92.5 98.7 100 80 225 1,760 180 1,800 Minor make-up

Source: Schendel et al. (1983)

four. When installation and maintenance costs are factored in, the two processes are considered to be competitive.

Oxo-alcoholSynthesis Gas In the production of oxo-alcohols, carbon monoxide is reacted with hydrogen at a one-toone ratio to form an aldehyde. The aldehyde is then reacted with pure hydrogen to form the desired oxo-alcohol product. The hydrogen and carbon monoxide synthesis gas is made by steam reforming of natural gas or by partial oxidation of hydrocarbons. The raw synthesis gas has a hydrogen to carbon monoxide ratio range of 3:1 to 2:1. Before the synthesis gas can be used, the ratio of hydrogen to carbon monoxide must be adjusted. Cryogenics, molecular-sieves, pressure swing adsorption (PSA), and membranes are used to make the ratio adjustment. The following case histories indicate how membranes have been used to increase capacity and/or efficiency in achieving the correct ratio of hydrogen/carbon monoxide for synthesis.

Membrane Permeation Processes

1261

MacLean and Graham (1980) reported on a membrane system used to debottleneck Monsanto's Texas City plant where the cold box did not have enough capacity to supply carbon monoxide feedstock for acetic acid production. The cold box had three product streams, a pure carbon monoxide stream for acetic acid, a 1.3:1 hydrogen/carbon monoxide stream for oxo-alcohols, and a hydrogen stream for methanol. The solution was to install a membrane separation system in parallel with the cold box. The 3.1:1 H2/CO feed was split between the membrane unit and the cold box. The cold box was then operated to produce hydrogen and the pure carbon monoxide required for acetic acid production. The membrane unit produced the 1.3:1 Hz/CO ratio stream required for oxo-alcohol production while recovering 93% of the carbon monoxide. The permeate stream, consisting of 96% pure hydrogen, was fed to the methanol plant. As an indication of membrane aging, the separation efficiency declined less than 10% during the first three years of operation. A membrane unit was also installed on the Texas City methanol plant purge gas stream (Burmaster and Carter, 1983). The membrane unit recovered about one half of the hydrogen and carbon dioxide that was previously sent to the fuel gas system. The methanol plant capacity was increased by 2.6%, and except for compressor limitations, could have been increased by 3.9%. The integration of a membrane system with pressure swing adsorption (PSA) at an aldehyde synthesis plant was studied by Doshi et al. (1989). Synthesis gas ratio adjustment using PSA alone produces high purity hydrogen as well as the proper syngas ratio. However, the drawback of the system is the relatively expensive tail gas compression that is required. PSA tail gas is produced at low pressure and requires a large compressor to boost the pressure to the aldehyde synthesis pressure. The use of a membrane system alone will produce syngas at the proper hydrogen/carbon monoxide ratio, but the purity of the hydrogen permeate stream is low. The syngas produced is at approximately the same pressure as the feed gas and can be used for aldehyde synthesis without additional compression. The hydrogen permeate stream generally has to be upgraded, which usually results in the loss of carbon monoxide from the system. An integrated system that uses both membranes and PSA on the product gas from a partial oxidation (POX) unit, minimizes the drawbacks of either system when it is used alone. In the integrated system, Figure 15-11, the feed gas enters the membrane unit where the residue stream, enriched in carbon monoxide, is produced at approximately feed pressure. The lower pressure hydrogen permeate stream is fed to a PSA unit where the carbon monoxide is adsorbed and a high purity hydrogen stream is produced. The low-volume, CO-rich PSA tailgas stream is compressed and combined with the residue stream from the membrane unit. The study shows the integrated system to have lower capital and operating costs than either of the individual systems and to provide 100% recovery of hydrogen and carbon monoxide. An economic comparison of a stand-alone PSA and an integrated membrane-PSA system is presented in Table 15-6.

Catalytic Reforming Catalytic reforming is a process in which hydrocarbon molecules are structurally rearranged to higher octane forms. The reforming process is a net producer of hydrogen that, if recovered, can be used in hydroprocessing. The following case history of a demonstration membrane system for the recovery of hydrogen from a catalytic reformer unit (CRU) off-gas was described by Yamashiro et al. (1985).

1262

Gas Purification

Natural Gas + Oxygen + Steam

Recycle ='- . . . . Ir (Optional)

CO= T

Unoptimized H=:CO Ratio / Lr ~

Syngas H=:CO= 1:1 Enriched CO

Enriched H= Product

Figure 15-11. Integrated membrane plus PSA system in a natural gas partial oxidation (POX) plant. (Doshi et al., 1989)

Table 15-6 Comparative Economics of PSA Alone vs. a Membrane/PSA Integrated System for Producing Hydrogen and Aldehyde Syngas Basis: Raw Feed Rate, MMscfd Synthesis Gas Rate, MMscfd Hydrogen Product Rate, MMscfd Raw H2/CO Gas Pressure, psig Aldehyde Reaction Pressure, psig Hydrogen Use Pressure, psig COMPOSITIONS, MOL- % H2 CO Ar/N 2 CH 4 CO2 H20

RAW GAS H2/CO 63.4 35.4 0.4 0.8 trace Sat'd

Compression Required, BHP Separation Equipment Cost, $MM US Installation Cost, $MM US Installed Compressor Cost, $MM US Compressor Operating Cost, $MM US (3 yr-8000 hr/yr-5r Total Capital Cost + 3 yrs Operation, ($MM US) Source." Doshi et al. (1989)

20.0 14.4 5.6 420 400 150 H2 PRODUCT 99.999 10ppm trace trace trace Dry

SYN. GAS 49.1 49.1 0.6 1.2 trace Sat'd

PSA ONLY 1080 1.425 0.175 0.864

MEMBRANE + PSA 415 1.375 0.225 0.332

0.966 3.430

0.371 2.303

Membrane Permeation Processes

1263

The 850 Mscfd unit was installed in the No. 2 Reformer at the Cosmo Oil Refinery in Chiba, Japan. The process flow for the unit is shown in Figure 15-12. The reformer effluent is cooled, the liquid and gas separated, and the gas fed to an absorber to remove the heavy components. Gas from the absorber, containing approximately 80% hydrogen and 20% methane and saturated with absorption oil components at a dew point of 95~ and a pressure of 398 psia, is fed to a filter separator to remove any residual liquids. It is then preheated to a temperature above the hydrocarbon dew point of the residual gas leaving the membrane unit and fed to the membrane separator. Heating the gas prevents condensation of the heavy hydrocarbons as the gas dew point increases with hydrogen removal, The 97% hydrogen permeate gas is of sufficient purity and pressure that it can be fed directly into the 256 psia hydrogen supply system. Typical operating conditions are provided in Table 15-7. The unit was operated at a low pressure ratio to deliver hydrogen to the refinery hydrogen supply system without recompression. This low pressure ratio resulted in a low hydrogen recovery of about 30%. During startup, the temperature of the feed was varied from 104 ~ to 180~ At higher temperatures the gas permeation rate increased by more than a factor of two. A corresponding decrease of membrane selectivity was also noted, but was not great enough to alter the system performance significantly. The permeate gas pressure was also varied while maintaining constant feed pressure. It was shown that membrane selectivity and hydrogen permeation rate are independent of differential pressure and pressure ratio for the range of conditions studied (Schell and Houston, 1985).

Full PI1BIl(I~6r

~ --

F'd~ersepatat~ P,ecyc~

Re~

Toturth=r

Figure 15-12. Flow diagram of reformer system with membrane hydrogen recovery unit on offgas ( Yamashiro et a/., 1985). Reproduced with permission from Hydrocarbon

Processing, February 1985

1264

Gas Purification

Table 15-7 Operating Conditionsfor Catalytic Reformer Offgas Hydrogen RecoveryMembrane System Feed Flow, Mscfd Composition, vol %: H2 CH 4 Dew Point,~ Inlet Pressure, psig Inlet Temperature,~ Hydrogen Recovery, %

Permeate

Residue

850 80 20 95 380 140 30

97 3

74 26

260

Source: Yamashiro et al., 1985

Another example of membranes being used for the recovery of hydrogen from a CRU was reported by Lane (1983). The unit feed was the offgas from a CRU that formerly went to a hydrogen plant. The feed gas flow rate was nominally 4 MMscfd, but varied from 2 to 10 MMscfd. The membrane unit typically produced 98% hydrogen at 250 psig with 36% recovery. The system operated over full CRU cycles, from start-of-run to end-of-run, which caused the hydrogen content of the feed to range from 62 to 87%. The membrane unit percent hydrogen recovery and product hydrogen purity remained relatively constant. To lessen scaleup risks, the smallest commercial-size membrane units were installed even though they were undersized for the flow. This resulted in the low 36% hydrogen recovery. The predicted recovery for the proper size membrane system is 93% with a hydrogen purity of 94%.

Hydroprocessing Hydrotreating, hydrodesulfurization, and hydrocracking are operations in which hydrogen is used to saturate the olefins in a hydrocarbon stream; remove objectionable elements such as sulfur, nitrogen, oxygen, halides, and trace metals; and crack larger hydrocarbon molecules into smaller ones. In these processes, fresh hydrogen is fed to the reactor, and unconsumed hydrogen is separated from the reactor effluent and recycled back to the reactor. A portion of the recycled hydrogen is often purged from the system to prevent the build up of light hydrocarbons and inerts that would lower the hydrogen partial pressure in the reactor. H y d r o t r e a t i n g . At Conoco's Ponca City Oklahoma refinery, a 71 mole % hydrogen highpressure.purge gas stream from a gas-oil hydrotreater was split to feed the light-cycle oil hydrodesulfurizer and the cryogenic liquified petroleum gas (LPG) recovery unit. The purge stream was used on a once-through basis in both units and then discharged to the fuel gas system. The installation of a membrane-based hydrogen recovery unit to produce high purity hydrogen from this purge stream was described by Shaver et al. (1991). A schematic of the system is presented in Figure 15-13. The membrane unit is designed to produce 7 MMscfd of 95 mole % hydrogen from the 12 MMscfd hydrotreater high pressure purge with a 75%

M e m b r a n e Permeation Processes

1265

Hydrogen recycle Membrane system ~ ~Y _~J.~..~--~GOHDT

Hydrogen I

J,o

I Makeup ! compressor

Fuel gas Cryogenic [

un.J

Figure 15-13. Flow diagram of membrane system installed to purify gas oil hydrotreater (GOHDT)offgas-to provide hydrogen for light-cycle oil hydrodesulfurizer (LCOHDS) (Shaver et al., 1991). Reproduced with permission from Hydrocarbon Processing, June 1991 hydrogen recovery. The high-purity permeate hydrogen is sent to the hydrodesulfurizer. The available pressure drop from 1,050 psig at the hydrotreater to 430 psig at the hydrodesulfurizer provides the driving force for the membrane separation. Because of the high purity of this permeate hydrogen, the hydrodesulfurizer offgas stream is still 90 plus percent hydrogen, and can be recycled back to the gas oil hydrotreater. The residue stream from the membrane unit is fed to the cryogenic unit for recovery of LPG. The pretreatment of the feed to the membrane unit consists of a knock out drum, a feed preheater, and a dry gas filter. The economics of this system is presented in Table 15-8 and indicates a 1.7-year payback period.

Table 15-8 Economics of Membrane Hydrogen Recovery System on Hydrotreater Offgas Investment, $ Debits, $/yr Steam consumption Lost hydrogen fuel value Maintenance and overhead Total Credits, $/yr Gas oil HDT product upgrade LCO HDS product upgrade Reduced power consumption Total Earnings, $/yr Simple payback period, yr Source: Shaver et al. (1991)

662,000 22,000 75,000 29,000 126,000 396,000 74,000 42,000 512,000 386,000 1.7

1266

Gas Purification

Operating data collected over two years show consistent recovery rates and permeate purity even with feed rate variation from 8 to 20 MMscfd. Schendel et al. (1983) compared membrane and PSA technologies for hydrogen recovery from hydrotreater purge gas. A 7 MMscfd purge stream at 800 psig and 100~ with 72% hydrogen was assumed to be the feed to the hydrogen recovery units. The product hydrogen could be returned to either the 250 psig make-up compressor first stage suction or to the 450 psig inter-stage suction. The residue gas is used for fuel gas. For the membrane separator, a 93% hydrogen product is produced at 250 psig with an 81% hydrogen recovery. The residue gas is let down in pressure and fed into the 60 psig fuel gas system. The feed gas to the membrane unit is preheated to avoid condensation. Figure 15-14 shows the integration of a hydrogen recovery unit into a typical hydrotreater unit with the recovered hydrogen returned to the suction of the low-pressure hydrogen makeup compressor. Pressure swing adsorption (PSA) utilizes molecular sieves to selectively remove hydrocarbons and other impurities to produce a high purity hydrogen stream. The greater the pressure swing (between the high pressure of adsorption and the low pressure of desorption), the greater the unit capacity and product recovery. The PSA system evaluation was based on two different operating scenarios: product at 450 psig and residue gas at 60 psig, and product at 250 psig and residue gas at 5 psig. In the 450 psig product case, the hydrogen is returned to the compressor interstage suction and the residue gas is fed to the 60 psig fuel gas system. In the 250 psig product case, the hydrogen is returned to the 250 psig compressor suction and the residue gas is sent to a low pressure burner. The hydrogen recovery was much greater for the low pressure case. The operating conditions and the operating and capital costs for the three cases are presented in Table 15-9. As can be seen from the cost data, the membrane system shows a

Residue

. H2 Make-up

Hydrocarbon Feed "

"'"ec~176

""~~

=,= ~al

~~

Make-up

Compm~o~ 7

Hydrotreating Reactor

i ~I

Recycle

Compressor 1 T

.P

9= Low Pressure v Purge

S.p, to,) I

'~

To Fractionation

Figure 15-14. Integration of hydrogen recovery unit into typical hydrotreater system. ($chende/ et al., 1983)

M e m b r a n e Permeation Processes

1267

Table 15-9 Operating Conditionsand Economicsfor HydrogenRecoveryfrom Hydrotreater Purge Gas by Membrane and PSA Systems

Waste Gas Pressure, psig Feed H2, % Feed Pressure, psig Feed Temperature,~ Feed Flow Rate, MMscfd High Purity H 2, % High Purity H 2, psig High Purity H 2, MMscfd Waste Gas H2, % Waste Gas Flow Rate, MMscfd H2 Recovery, % Capital Costs, M$ Equipment Installation Total Cost Operating Costs, M$/yr (2) Steam for Preheat Compression to Reactor Press. Contrib. to H2 Cost, $/Mscf Total H 2 Cost, $/Mscf (3)

Membrane

PSA

PSA

(1) 72 800 100 7 93 250 4.42 36 2.58 81

60 72 800 100 7 99.5 450 3.04 51 3.96 60

5 72 800 100 7 99.9 250 4.05 34 2.95 80

530 100 630

1,050 175 1,225

875 150 1,025

10 140 0.11 0.20

46 0.05 0.29

w 130 0.09 0.24

Notes: 1. Membrane system residue gas produced at operating pressure and, after pressure reduction, fed into the fuel gas system. 2. Utili~ costs based on 5c/kwh and $5/MM Btu. 3. He cost based on 8,000 hr/yr operation for a five-year span. Source: Schendel et al. (1983)

lower capital cost and lower total hydrogen cost than either of the PSA cases. The membrane and the low-pressure PSA case incur a significant cost for recompressing the hydrogen to reactor pressure. The low-pressure PSA case has the better economics of the two PSA cases. However, it may be difficult to find a use for the 5 psig waste gas. Also, depending on the sensitivity of the hydrotreater to hydrogen partial pressure, the higher purity of the PSA gas may have definite advantages. Tonen Technology K.K. and UBE Industries (1990) reported on the installation of a hydrogen recovery facility at the Wakayama refinery of Tonen Company, Ltd. Polyimide membrane modules, arranged in eight rows of two trains each, were designed to produce 5,153 scfm of product hydrogen with a minimum purity of 95%. The feed is filtered to remove hydrocarbon mist and then preheated to 160~176 prior to entering the modules. The hydrogen-rich permeate is collected and fed to the hydrogen system. The residue gas is collected, cooled to 150~ and discharged to the fuel gas system. Plant operating data are presented in Table 15-10.

1268

Gas Purification

Table 15-10 Operating Data for HydrogenRecoveryMembrane System at Wakayama Refinery Feed Flow Rate, scfm Pressure, psig Temperature,~ Composition, mol % H2 CI C2 C3 C4+ H2S, ppm BTX, ppm Mw H2 Recovery, %

15,490 343 90 77.6 17.0 4.4 0.3 0.7 400 300 6.30

Product

Offgas

10,095 124 155

5,395 102 99

98.2 1.6 0.2

300 2.3 82.5

39.1 45.8 12.2 .09 20 590 13.81

Source: Tonen Technology and UBE Industries (1990)

Hydrocracking. Hydrocrackers typically operate at higher pressures than hydrotreaters or hydrodesulfurization (HDS) units. Bollinger et al. (1984) performed a study to optimize hydrogen recovery from hydrocracker purge gas streams. Various membrane separation operating options were studied. Operating options included constant recycle purity, constant purge rate, constant make-up compressor horsepower, and constant hydrogen make-up rate. The optimized system, which includes the recovery of hydrogen from both the high- and low-pressure purge streams, is shown in Figure 15-15, which includes material balance data for the system. As indicated in the material balance, hydrogen is consumed during the hydrocracking reaction (chemical hydrogen) and some of the unreacted hydrogen is discharged in purge streams from the high- and low-pressure separators. The optimized design depicted in Figure 15-15 recovers hydrogen from the two purge streams thereby minimizing hydrogen losses. When compared to the hydrocracker without membrane units, the optimized system results in a 10% increase in the hydrogen partial pressure of the recycle gas leaving the highpressure separator, a slight reduction in make-up hydrogen, and an increased chemical hydrogen consumption (26,7 to 31.8 MMscfd). Assuming that chemical hydrogen consumption per barrel of feed remains constant, the increase in chemical hydrogen consumption corresponds to an increase in hydrocracker throughput of 19%. While the make-up hydrogen flow for the optimized system decreases from 40 to 38.7 MMscfd, the total flow through the make-up hydrogen compressor increases from 40 to 59.0 MMscfd as both permeate streams must be compressed in addition to the make-up hydrogen. " B u t a m e r " Offgas. In the UOP licensed "Butamer" process, normal butane is catalytically isomerized to isobutane. The process produces isobutane and hydrogen streams. Even with high hydrogen recycle rates, some feedstock is cracked into methane, ethane, and propane. To maintain high reactor efficiency, some of the recycle gas is purged, usually to the fuel gas

Membrane Permeation Processes

1269

Membrane Separators

High Pressure Membrane Separators Reactor

Stream Pressure, psig Flowrate, MMSCFD

(~ 250 38.7

(~) 1800 18.9

(~) 270 10.2

~ 1450 4.4

(~ 450 14.5

(~ 220 4.4

(~) 50 5.8

Composition, mole% Ha Cl Ca C3+

88.5 7.0 3.1 1.4

82.0 12.2 4.6 1.2

61.9 21.1 10.0 7.0

34.8 43.3 17.1 4.8

96.5 2.6 0.7 0.2

21.8 42.3 20.9 15.0

91.7 5.3 1.8 1.2

Figure 15-15. Optimized flow scheme for recovery of hydrogen from hydrocracker purge streams (Bollinger et a/., 1984). Reproduced with permission from Chemical

Engineering Progress, copyright 1984, American Institute of Chemical Engineers

system. This stream contains not only hydrogen and the cracking products, but also some butanes. Membranes have been considered for the recovery of hydrogen from Butamer units. However, to maintain catalyst activity, small amounts of an organic chloride are introduced into the feed stream. The chlorides are converted to HC1 by the catalyst, and the purge gas from the Butamer unit contains traces of HC1. The effect of HC1 on membranes is a concern. A simplified process schematic for a "Butamer" unit with hydrogen recovery is shown in Figure 15-16. The membrane unit is located downstream of a caustic wash used to remove HC1 from the purge gas. Schell and Houston (1982) described the integration of a membrane unit with a "Butamer" unit, designed to process 47,800 scfh of feed gas. The flow rates and purities are shown in Table 15-11. They report that under bone-dry feed conditions, the cellulose acetate membrane, which was located upstream of the caustic wash in this plant, was not affected by HC1; however, special materials and adhesives were used to ensure resistance to the HC1. Cooley and Dethloff (1985) reported on a demonstration unit installed initially upstream of the caustic wash unit. They found that HC1 concentrations of 2,000 and 4,000 ppm in the purge gas impaired the membrane performance. The unit was relocated downstream of the caustic wash as shown in Figure 15-16 and the performance improved. The unit was designed for a feed rate of 971 scfh at 295 psig with 61% hydrogen in the feed gas.

1270

Gas Purification

Table 15-11 Operating Data for HydrogenRecoveryfrom the Butamer Process

Pressure, psia Flow Rate, Mscfh Temperature,~ Composition, mol % H2 C1 C2+ HC1

Feed Gas

Residual Gas

Permeate Gas

265 47.8 110

240 16.8 100

15 31.0 100

68.9 23.7 6.8 0.6

17.8 63.0 19.0 0.2

96.4 2.6 0.2 0.8

Source." Schell and Houston (1982)

Recovered Hydrogen

Make-up ( Hydrogen J

Recycle "

/ Caustic Wash N-Butane

Butamer Unit

/

Membrane Unit

L oc,.

Recovery

Isobutane Product

Figure 15-16. Membranesystemfor recoveringhydrogenfrom Butamersystempurge gas. (Cooleyand Oethloff, 1985)

OtherHydrogenApplications MacLean and Narayan (1982) have described other applications of membrane systems for hydrogen separation, including toluene hydrodealkylation and coal liquefaction processes.

Carbon Dioxide, Hydrogen Sulfide, and Water Removal Enhanced Oil Recovery (EOR) When CO2 is injected into an oil reservoir at sufficient pressure, it dissolves in the oil present in the substrata reducing its viscosity, allowing it to flow more freely, and thereby increasing oil production. When the oil is brought to the surface and its pressure reduced, the injected CO2 is released from the oil and discharged with the associated gas.

Membrane Permeation Processes

1271

In C O 2 flood EOR projects, the C O 2 is typically recovered from the associated gas and recycled back into the oil producing formation. With this continual recycle of CO2, both the volume and CO2 content of the associated gas progressively increase. In the design of CO2 flood EOR systems, the objective is to minimize capital expenditure when the associated gas volume and CO2 content are low, but have enough design flexibility so that the system will be operable in the future when the associated gas volume and CO2 content are high. In many ways, the modular nature of membranes makes them ideally suited for this application. In the initial phases of the EOR project, capital costs can be minimized by adding the minimum number of membrane modules and CO2 re-injection compressors. Additional membrane modules and CO2 compression can be added later when the associated gas volume and CO2 content are higher. Other CO2 recovery processes do not have this flexibility. The first commercial scale membrane installation used with CO2 flood EOR was the Sacroc project in West Texas. Injection of CO2 into the field at volumes up to 200 MMscfd began in 1972. To handle the anticipated increase in the associated gas CO2 concentration, Sacroc installed three CO2-removal facilities. The plants were installed in conjunction with three existing processing facilities operated by Sun Exploration & Production Co., Chevron U.S.A., and Monsanto. The Sun and Chevron facilities use the Benfield hot potassium carbonate process, and the Monsanto facility uses the monoethanolamine (MEA) process. The Sun hot potassium carbonate plant was designed to reduce the CO2 content of 160 MMscfd of associated gas from 24 to 0.5% CO2, while the Chevron plant was designed to reduce the CO2 concentration of 46 MMscfd of associated gas from 24 to 1.0% CO2. The Monsanto MEA plant was designed to treat 16.5 MMscfd of the 24% CO2 feed gas, reducing its CO2 concentration to 0.01% (Parro, 1984). In the late 1970s, Sacroc realized that the CO2 content of the gas produced from the field would peak for a few years at levels greater than the CO2-removal plants' capacity. The CO2 content of the field gas had gradually increased from 0.5% prior to injection to 40% CO2. Sacroc contracted with The Cynara Company to build and operate two membrane units. The new membrane units were installed and operated integrally with the Sun and Chevron hot potassium carbonate plants. The Sun membrane unit was designed to recover 50 MMscfd of CO2 at 520 psig with an allowable pressure drop of 40 psig. The Chevron membrane unit was designed to recover 20 MMscfd of CO2 at 480 psig with an allowable pressure drop of 40 psig. The membrane separation process flow is given in Figure 15-17. The two important features of this design are the dehydration of the inlet gas and the operation of the membranes at reduced temperatures. The inlet gas is first cooled by cross exchange with the CO2 and hydrocarbon product gas streams. This reduces both the moisture and hydrocarbon content of the feed gas and the size of the dehydration equipment. After dehydration, the feed gas is cooled by cross exchange with the residue gas and finally refrigerated before being directed to the membrane modules. The low-temperature feed results in a higher separation factor and a reduced volume of gas. The membranes are single stage and are arranged in parallel. The amount of feed gas processed depends on the field production rate, and changes in flow were handled by increasing or decreasing the number of membrane modules on line. When the plant was shut down, depressurized, and restarted; the permeate gas flow was found to be 5 to 15% lower than previous levels. The permeate flow gradually improved over a few days of operation, but did not recover entirely. Some loss of flux remained (with a slightly better separation factor), and additional modules were required to maintain the same permeate flow rate (Marquez and Hamaker 1986).

1272

Gas Purification

Acid Gas

Cross Exchange Acid.Gas

I

Hydrocarbons,Gas "~

Dehydration

Cross Exchange,

and Refrigeration t

I

" ]

I

"

=~epa~ "

-

HydrocarbonGas (Residue)

Figure 15-17. Membrane system for treating gas at Sacroc enhanced oil recovery (EOR) project in West Texas. (Cutler and Johnson, 1985) Additional field test data on the use of membrane permeation to provide a concentrated CO2 stream for EOR were reported by Russell (1984) and Mazur and Chan (1982). Estimated capital costs for processing a 800 psig, 20 MMscfd feed gas stream containing varying amounts of CO2 were developed by Coady and Davis (1982) and are presented in Figure 1518. In all cases, the permeate stream contains 5% hydrocarbons while the residue gas contains 2% CO2. To obtain a 95% CO2 permeate stream, a two-stage membrane system with interstage recompression is required for feed gas CO2 concentrations less than 75 vol%. At feed gas CO2 concentrations greater than 75%, the second-stage membrane and interstage recompression can be eliminated. Low permeate pressure results in the least amount of membrane area required to achieve the desired separation. However, this is at the expense of recompression horsepower required for reinjection. Marquez (1991) reported on a series of tests performed to determine membrane performance with high-permeate back pressures. The tests covered feed gas pressures from 313 to 363 psig (333 psig average) and permeate back pressures from 189 to 250 psig (221 psig average). The permeate composition averaged about 97% CO2 and 0.13% H2S, representing removal of about 37% of the CO2 and 40% of the HzS from the feed gas. It was concluded that permeation into a high back pressure system is feasible and can result is considerable recompression cost savings. New membrane modules were used for the tests. Permeation rate measurements showed that the flux stabilized after a month of operation at 79% of the initial rate. In work sponsored by the U.S. Department of Energy (1989), the cost was developed for a membrane unit installed upstream of an amine unit. The membrane unit was designed to process 170 MMscfd of feed gas containing 17% HzS and 45% CO2. It was estimated that 280,000 ft 2 of membrane would be required to remove 70% of the acid gas. At an installed first cost of $20 per ft 2 of membrane, the cost of the unit would be $5.6 million. The estimated annual steam savings in the amine plant were $5 million to $10 million based on 0.8 to 1.6 pounds of steam per pound of acid gas removed and a cost of $2.28 per 1,000 pounds of steam. The net annual savings, including membrane replacement, were $4.4 million to $9.4 million. Goddin (1982) compared several methods for recovering CO2 from a CO2-flood project associated gas stream. In this study, the associated gas hydrocarbon and nitrogen flow rates were held constant while the CO2 content increased up to 90 vol%. This simulates the change in associated gas composition over the EOR project life. The following CO2 recovery cases were evaluated:

Membrane Permeation Processes I

I

I

I

I

5 A

Feed Flow Rate = 20 MMSCF/D

I

1273

I

Compression Eliminated I

o

-- 4 :E m

v

if)

~3 o

O ii

Total Facility n~ (Includes Dehydratlo

m

/

== 2 Q.

I

~ ' - ~

Recompresslon I TotaIFacil

O 1 Membrane 10

20

30

40

50

60

70

80

90

CO= In Feed, %

Figure 15-18. Estimated capital costs for carbon dioxide recovery from a 20 MMscf/d gas stream ( Coady and Davis, 1982). Reproduced with permission from Chemica/

Engineering Progress, copyright 1982, American Institute of Chemica/ Engineers

1) Conventional Amine The amine unit is based on a 30% DEA solution. The feed gas is compressed to 285 psia and pretreated to remove heavy hydrocarbons. After passing through the DEA absorber, the sweet offgas is compressed to 650 psia and sent to the existing gas plant. The sour CO2 stream from the DEA stripper, at 20 psig, is compressed to 450 psia and sent to a Selexol sweetening process designed to reduce the HzS content of the CO2 product gas to less than 100 ppm. 2) Cryogenic Fractionation (Ryan-Holmes Process) The feed to the cryogenic unit is compressed to 625 psia, dehydrated, and chilled. The overhead gas from the first cryogenic column, which contains the methane and lighter components, is sweet pipeline gas. The bottom product from this column flows to the CO2 column, where the sweet CO2 leaves in the overhead. Lean oil is added to the top of the first column to prevent the CO2 from freezing, and also is added to the top of the second column to break up a CO2/ethane azeotrope. The bottoms from the CO2 column are sent to a lean oil recovery unit. The propane and lighter components are processed in a DEA unit to remove the CO2 and HzS. The acid gases removed here are processed in a Claus unit for sulfur recovery. Two cases were considered. Case A assumes that all of the ethane recovered has the value of liquid hydrocarbon. Case B assumes that only a portion of the ethane has the high liquid hydrocarbon value since 80% of the hydrocarbons would be recovered in the existing gasoline plant. Credit was given for the differential value of ethane in natural gas liquids (NGL) versus fuel gas. 3) TEA Bulk C02 Removal In the TEA bulk removal process, a TEA solution is used to remove the H2S and CO2 from the feed. The CO2 and HzS are then removed from the TEA solution by flashing it to about 20 psia. The TEA absorber overhead stream, containing about 20% CO2 and some HzS, is sent to a DEA unit for final clean up. The acid gas streams from the DEA unit and the TEA flash tower are compressed to 450 psia and sent to a Selexol unit for HzS removal.

1274

Gas Purification

4) Membrane Permeation The membrane unit was designed to produce a permeate with a maximum of 5% hydrocarbons and a residue stream containing 20% CO2. The residue stream is sent to a DEA unit for cleanup. The acid gas from the DEA unit and the sour permeate streams are compressed and sent to a Selexol unit for H2S removal. To develop a range of costs, two cases were considered. The low cost case assumed a short membrane module life and a high permeation rate. The high cost case assumed a long module life with a lower permeation rate. A summary of capital costs versus feed rate is given in Table 15-12. Figures 15-19 and 15-20 present the cost of removing CO2 for two different sets of utility costs. Also plotted are the cost curves for sweetening the CO2 product stream using the Selexol process. The cost of the Selexol process is included in the curves for the DEA, TEA, and membrane processes. For the low energy cost case (Figure 15-19), the least cost system, over most of the CO2 concentration range, is the cryogenic process with full credit for ethane recovery. When only partial credit is taken for ethane recovery, the cryogenic system is not competitive. Membrane permeation is more economical than DEA and TEA over the entire range, and has an increasing advantage over DEA at CO2 concentrations above 20%. The effect of higher energy cost (Figure 15-20) is to make TEA and the membrane process more economical than any of the others over the entire CO2 concentration range. In the mid 1980s, several large CO2-flood projects were initiated based on the availability of naturally occurring CO2 brought in by pipeline to West Texas. At that time crude prices were high and, significantly, NGL prices were also high. Overall economics were quite similar to the scenario presented by Goddin in Figure 15-19. Three of the large projects: Amerada Hess's Seminole Plant (2 x 85 MMscfd, 77% CO2) (Schaffert et al., 1986; Wood et al., 1986), Shell's Wasson Denver unit (275 MMscfd, 93% CO2) (Flynn, 1983; Youn et al., 1987), and Arco's Willard unit (72.9 MMscfd, 86% CO2) (Price and Gregg, 1983), and a smaller plant, the Mitchell Alvord South CO2 plant (7.5 MMscfd, 85% CO2) (McCann et al., 1987) used cryogenic distillation (the Ryan Holmes process). All of these projects recovered NGL and cryogenic distillation was chosen because economics favored NGL production.

Table 15-12 Summary of C02 Removal Costs for DEA, Cryogenic,TEA, and Membrane Facilities Percent C O 2 Flow, MMscfd Capital, $MM DEA Cryo A Cryo B TEA-DEA Perm-Low Perm-High Source: Goddin (1982)

20 18.2

40 24.3

60 36.5

80 73.3

90 148.0

9.7 16.2 16.2 --~

16.4 20.8 20.8 15.0 13.8 16.6

25.5 28.3 28.3 21.6 18.8 23.4

54.5 42.6 42.6 36.9 29.8 37.9

103.6 73.5 73.5 65.0 47.0 60.5

Membrane Permeation Processes

1275

FUEL GAS @$2.001MM BTU NGL @$5.001MM BTU POWER @ $0.051KWH

CRYO-B (',,,m 0

=.3 0

~E

"2 o.,

DEA PERM

CR

~ C O 2 SWEETINING 0 0 co 2 iN F E D - ~ L

Figure 15-19. Estimated costs of carbon dioxide removal for a C02-flood EOR project-low utility cost case (Goddin, 1982). Reproduced with permission from Proceedingsof the 61st Annua/ Conventionof the GPA, copyright 1982, Gas ProcessorsAssociation

\~\~

FUEL G A S @ S6.001MM BTU NGL @ $7.001MMBTU WH

4 u,. O u.

3

=E r162 MJ o.

"

CRYO

2

\-~ TEA ~ l ~~/CO

l 2 SWEETENING

PERM

CO2 IN FEED - MOL %

Figure 15-20. Estimated costs of carbon dioxide removal for a C02-flood EOR project--high utility cost case (Goddin, 1982). Reproduced with permission from the 61st Annual Convention of the GPA, copyright 1982, Gas ProcessorsAssociation

1276

Gas Purification

Chevron's Sacroc project (Parro, 1984; Schendel and Nolley, 1984) and Amoco's Central Mallet unit (100 MMscfd, 85% CO2) (Anon., 1985) used membranes. However, the Sacroc project predates the Ryan Holmes process, and Amoco was under contractual obligation to supply gas with NGL's still present to a downstream NGL extraction plant. Therefore in the mid-1980s, the market generally sustained the conclusion of Goddin (1982) that, when NGL was recovered, economics favored cryogenic distillation over the use of either membranes or amine treating in CO2-flood EOR projects. Interest in CO2-flood EOR fell with the drop in oil prices that occurred in the early 1980s, but seems to be reviving. Recent CO2-flood EOR projects using membranes include Mobil's Salt Creek project (64-110 MMscfd, 70% CO2), which started up in 1992; Shell's McCarney, Texas plant, which began operation in 1993 (11 MMscfd, 70% CO2); and Amoco's Mallet plant in West Texas (30-100 MMscfd, 80% CO2), which started up in 1994 (Cynara, 1995). One of the key factors favoring membranes in these more recent projects is the ability to delay capital expenditures. Schendel (1984) proposed the integration of a membrane process and cryogenic distillation. Feed gas from the field passes through a hydrocarbon dew point control unit and then through a membrane unit. This first membrane unit removes the bulk of the CO2. The residue is fed to the cryogenic unit. Rather than suppress the CO2/ethane azeotrope with a hydrocarbon recycle stream, the CO2-rich azeotrope is allowed to go overhead to a second membrane unit where CO2 is removed as the permeate. The primary advantage of this scheme is that the cryogenic unit can be sized for the relatively constant hydrocarbon rate, not the increasing CO2 rate over time. As the CO2 rate increases, additional front-end membrane modules can be added as required. The disadvantage of this concept is that the CO2 product is produced at low pressure, requiring recompression. Some cryogenic distillation process configurations can produce the CO2 product as a liquid which can be easily pumped to high pressure (Wood et al., 1986; Ryan and Schaffert, 1984). Boustany et al. (1983) compared the performance and economics of hot potassium carbonate, cryogenic separation, and membrane permeation processes for CO2 removal. They assumed a 100 MMscfd feed stream containing 80% CO2 available at 25 psig and 100~ The gas is assumed to be sweet. The cryogenic unit is similar to that previously described by Goddin (1982) except that the CO2 recovered from the DEA unit is compressed and combined with that coming from the cryogenic column overhead. The membrane unit consists of two sections in series. The permeate from the second section is compressed and recycled to the feed. The residue from the second stage is fed to a DEA unit. The CO2 recovered from the DEA unit is combined with that from the first membrane section and compressed to 500 psig. A summary of operating conditions and cost data is presented in Table 15-13. It is concluded that the membrane permeation units (PRISM separation) offer both capital and operating cost savings over the other systems.

Well Fracturing In this process, which is applicable to certain types of wells, high pressure C O 2 is injected into a reservoir to fracture the formation. A slurry of sand is then fed into the well to fill the fractures and provide a porous channel for gas or oil flow. The associated gas flow after the fracture contains a high amount of CO2, which decreases to pipeline transmission levels over a relatively short period of time. Membranes work well for treating the gas resulting from the fracturing process. Because of their modular nature and portability, they can be used to

M e m b r a n e Permeation Processes

1277

Table 15-13 Summary of Data for Comparison of Cryogenic, Hot Potassium Carbonate, and Membrane Processes

Stream Recovery, % CO2 Purity, % Hydrogen Recovery, % Capital Costs, $MM CO2 Removal DEA Treating Feed Compression 25 to 250 psig 250 to 500 psig Hydrocarbon Liquid Recovery CO2 Compression 10 to 500 psig Total Capital Operating Costs, $MM

Hot Potassium Carbonate

Cryogenic

Membrane

96.9 99.8 99.5

93.0 95.1 81.01

96.9 95.5 83.5

21.1 m

24.2 4.9

16.1 4.9

11.1 0.9 4.0 7.9 45.0 11.78

11.1 6.2

12.1 0.9

m 46.4 8.46

5.1 43.1 6.46

C O 2 Product CO 2

Notes: I. Assumes C02 column operation based on ethane rejection~propane recovery. Source: Boustany et al. (1983)

maintain pipeline quality gas until the C O 2 levels return to normal and can then be removed and used elsewhere (Schendel and Seymour, 1985).

Pipeline Natural Gas Typical specifications for pipeline gas are C O 2 less than 2%, H2S less than 4 ppmv, and water less than 7 lb/MMscf. Water vapor is a very fast gas and, therefore, dehydration to pipeline specifications can be achieved by membrane systems while also removing the slower CO2 or HzS acid gases. Membrane permeation systems can remove CO2 to the required 2% level; however, at low CO2 concentrations, the CO2 partial pressure driving force is reduced and a significant amount of hydrocarbon (primarily methane) is lost with the CO2rich permeate. The problem becomes even more severe when trying to remove substantial quantities of HzS to meet a 4 ppmv specification. Normally, with membranes only modest adjustments to HzS concentration in the ppm range are economically feasible. However, subsequent treatment with other processes to meet the H2S specification is possible. The Gas Research Institute (GRI) performed an extended field evaluation of a membrane unit operating on low quality natural gas (Meyer and Gamez, 1995; Lee et al., 1995). The unit was a standard two-tube Grace Membrane Systems (GMS) test system designed to treat 0.5 MMscf/d of 750 psig natural gas containing 6.0% carbon dioxide. Two types of asymmetric cellulose acetate membrane (standard and higher density) were tested. During the

1278

Gas Purification

573-day test period the unit operated smoothly, reducing the CO2 content of the gas to meet the pipeline specification of less than 2.0% carbon dioxide. The system also provided gas dehydration to 7 lb H20/MMscf or less. GRI evaluated process economics based on the test results. For the case of a plant treating 37 MMscf/d at 725 psig, the study indicated that the membrane process would be competitive with DEA or MDEA systems if the gas contained less than 20% CO2 and appreciably less expensive than the amine processes if the gas contained over 20% CO2 (Meyer and Gamez, 1995). Bhide and Stern (1993) conducted a study to optimize the process configuration and operating conditions and to assess the economics of membrane processes for removing CO2 from natural gas. The base case was a 35 MMscf/d natural gas feed stream at 800 psia with CO2 concentrations in the range of 5 to 40 mole%. The optimum configuration was found to be a three-stage system consisting of a single permeation stage in series with a two-stage permeation cascade with recycle. For the base case operating conditions, membrane properties, and economic parameters assumed in the study, membrane processes were found to be more economical than DEA over the entire range of CO2 concentrations considered (with no H2S in the feed). When H2S is also present, the results showed that the cost of meeting product gas specifications (<2 mole% CO2 and <4 ppmv H2S) increases with increased H2S concentration. For example, with feed containing 1,000 ppm H2S, the membrane process was found to be more economical than DEA only when the total acid gas in the feed exceeded about 16 mole%. Fournie and Agostini (1987) studied the removal of CO2 from gas that was to be sold as liquified natural gas (LNG) and from gas that was to be sold in the gaseous phase. For gas to be sold as LNG, with a 100 ppm CO2 specification, they concluded that membranes were 1.1 to 2.5 times more capital intensive than amine units. Three feed CO2 concentration ranges were investigated for gas to be sold as a gaseous product containing less than 2% CO2. For a feed concentration range of 5 to 20% CO2, membrane systems were estimated to be one-half the investment cost of diethanolamine processes, but were not recommended because of the low hydrocarbon recovery and the problem of finding a use for the low pressure permeate. The predicted investment cost for CO2 concentrations ranging from 20 to 40% was about the same as that for conventional processes, but the hydrocarbon losses were significant. For CO2 concentrations greater than 40%, the estimated investment was about one-half that of conventional processes. Their overall conclusion is that membranes are most effective for high CO2 concentrations, especially when the CO2-rich permeate can be used for EOR applications. The capital costs of membrane systems to remove H2S to meet a sales specification of 4 ppm were found to be about two times the cost of conventional systems. The use of membranes for bulk removal of H2S down to about 2% combined with conventional processing to bring the concentration down to 4 ppm resulted in capital costs about the same as conventional processes, but with a much lower overall energy consumption (Fournie and Agostini, 1987). The use of natural gas as the circulating fluid drilling gas in a deep gas formation (rather than air or conventional muds) was reported by Cooley and Dethloff (1985). The gas available for this use was saturated with water and contained 6% CO2 and 3,000 ppm H2S. The CO2 and H2S levels were considered too high so a gas purification system based on membrane permeation was installed. The membrane unit reduced the CO2 content to 2%. The H2S content was reduced to 600 ppm, which then could be economically treated with an iron sponge unit. Operating data are presented in Table 15-14. Some of the early tests demonstrating the feasibility of using membranes to remove CO2 from natural gas to produce a sales grade gas have been reported by Mazur and Chan (1982)

Membrane Permeation Processes

1279

Table 15-14 Operating Data for Drilling Gas Application of Membrane Process

Flow, MMscfd Pressure, psia Temperature,~ CO2 Content, mol % H2S Content, ppm

Inlet

Product Gas

1.22 900 100 6.05 3000

1.07 900 90 2.0 600

Source: Cooley and Dethloff (1985)

and Schell and Houston (1982). Pipeline gas with less than 5% C O 2 w a s produced from wellhead gas containing up to 30% CO2. Inlet pressures ranged from 250 psig to 800 psig. Another economic comparison of membranes and amine treatment was presented by Babcock et al. (1988). They considered CO2 feed gas levels ranging from 10 to 60 vol% as well as various operating pressures and flow rates. The base case feed gas contained less than 0.5% HzS and had a flow rate of either 37.2 or 60 MMscfd at 725 psig and 100~ The product gas specification called for CO2 and HzS concentrations to be less than 2% and 4 ppm, respectively. The conventional DEA and MDEA split-flow/flash amine systems operated at 900 psia. The circulating solutions were 30% DEA and 50% MDEA. Acid gas in both systems was released at 25 psia. The multistage membrane system consisted of an initial membrane bulk removal unit followed by a two-stage permeation system with recycle. Some of the estimated capital and operating costs are summarized in Table 15-15. Processing costs include operating expenses, lost product gas, and capital charges. The lost product gas includes gas lost in the permeate and gas used for fuel for recycle recompression. The processing cost comparison is presented graphically in Figure 15-21. The plotted results indicate that membrane processes are more economical than conventional DEA over the entire feed CO2 concentration range. They are more economical than two-stage MDEA at CO2 concentrations below about 15% and above about 42%; while the MDEA system is slightly more economical in the range between these two concentrations. McKee et al. (1991) prepared an economic comparison of membrane, amine, and membrane/amine hybrid processing systems to aid in determining when the hybrid system would have an economic advantage. In the hybrid system, the bulk of the CO2 is removed with a small area membrane system and the amine system is used to remove the residual CO2 to meet specification. Feed gas CO2 concentrations from 4 to 27% and flow rates between 5 MMscfd and 75 MMscfd were used to generate cost data. The amine system design parameters were 35% DEA, 70% approach to equilibrium, and a reboiler steam rate of 0.7 pounds per gallon of DEA circulated. The membrane was designed as a single stage with a 20 psig permeate pressure. Feed gas values of $0.50/MMBtu to $2.00/MMBtu were used. The cost comparisons at $1.50/MMBtu feed gas value show that 1) Amine systems are preferred when flow rates are high (over about 20-30 MMscfd) and CO2 concentrations are low (below about 16 mole%).

1280

Gas Purification

Table 15-15 Capital and Operating Costs for Amine and Membrane Systems C O 2 in Feed, %

10

20

Flow Rate, MMscfd DEA Capital, $MM Processing, $/Mscf Multistage Membrane Capital, $MM Processing, $/Mscf Single-Stage Membrane Processing, $/Mscf

37.2

37.2

Flow Rate, MMscfd MDEA Capital, $MM Processing, $/Mscf

60.0

25

30

50

80

37.2

37.2

37.2

4.54 0.24

6.21 0.36

7.50 0.46

9.50 0.64

11.80 0.87

3.33 0.19

3.69 0.25

3.37 0.27

2.45 0.24

1.44 0.14

0.26

0.32

0.32

0.28

7.32 0.21

60.0 7.60 0.23

0.15 60.0 8.76 0.30

Notes: 1. Feed gas conditions are 725 psig and lO0~ for all cases. All feed gases have less than 0.5 vol% H2S. 2. In all instances product specifications are C02 < 2 vol%, H2S < 4 ppm. Source: Babcock et al. (1988)

2) Membrane systems have an advantage at low feed rates (below about 20-30 MMscfd) over the entire range of feed gas compositions. 3) The hybrid system is favorable at high flow rates (above about 30 MMscfd) and high CO2 concentrations (above about 16 mole %). Membrane units have a smaller plot space requirement and weigh less than most other process units. These factors, as applied to an offshore installation, were evaluated by Cooley (1990). A comparison was made between an amine/glycol unit and a membrane unit designed to treat 96 MMscfd of natural gas at 1,000 psig. The systems were designed to reduce the CO2 content of the gas from 15.8 to <2% and dehydrate the gas to North Sea specifications. The results of the comparison are presented in Table 15-16.

Landfill Gas Purification Landfill gas is produced at low pressure (near atmospheric), has a high CO2 content, and contains numerous trace contaminants. The operation of a membrane system producing approximately 800 Mscfd of high Btu gas at 500 psig, containing less than 3.5% CO2 was described by Houston and Schell (1986). The feed gas from the collection system is first compressed to 525 psig. The gas is then treated in a carbon bed adsorber to remove organics, fed to a filter/separator to remove condensed liquids, and preheated to 130~ to prevent water condensation inside the membrane. The gas is then fed to a two stage membrane unit

Membrane Permeation Processes

$0.60 A V) t~

r

1281

/

$0.50

"O IL

,,o

$0.40

/

/

u)

AMINE PROCESSING Amine

O, /

$0.30

MDEA "two Stage . . _ ._-- - l - - " ......._._ ..........- ------ " -

/

"-" -

$0.20

r

$0.10

t~t~g ''~

MEMBRANE PROCESSING

$0.00

I

I

[

I

I

i

I

i

l

0% 10% 20% 30% 40% 50% 60% 70% 80% 90% 100% PERCENT C02 IN FEED (%)

Figure 15-21. Cost comparison between amine and membrane processes for purification of 60 MMscf/d of natural gas (Babcock et a/., 1988). Reproduced with

permission from Energy Progress

Table 15-16 Plot, Cost, and Weight Requirements for Membrane and Amine/Glycol Dehydration Acid Gas Removal Units

Weight, metric tons Installed Cost, $MM Deck Area, sq. meters

Amine Plant

Membrane Plant

1,605 30.5 600

503 18.7 214

Source: Cooley (1990)

with recycle. The CO2-rich permeate from the first stage is vented and the second-stage permeate is compressed and recycled to the first stage to improve methane recovery. Advantages claimed for membrane systems in this application are low capital and operating costs, simplicity of operation and maintenance, compact size, and modular construction. The unit also dehydrates the gas.

Helium Removal from Natural Gas Helium occurs naturally as a minor constituent (<5%) of some natural gas streams. It is usually recovered by a two-step process: the first recovers a crude helium concentrate (about 50%) from the natural gas, and the second purifies the helium to a high purity product (>99.99% for Grade A Helium). Several hybrid processes for accomplishing both steps of the helium recovery and purification operation have been evaluated by Choe et al. (1988).

1282

Gas Purification

The most economically attractive process is a hybrid system that combines cryogenic, PSA, and membrane units to produce 99.99% helium from dilute natural gas. The process is shown in Figure 15-22. The feed is natural gas containing 2.1% helium. The feed is first cooled to -60~ to condense the heavier hydrocarbons. Then the gas is cooled to -240~ to condense most of the methane and some nitrogen. At this point the gas contains 30-35% helium. The crude helium is then fed to a two-stage membrane unit that produces a 95% helium stream. PSA is used to upgrade this stream to Grade A purity.

Air Separation The initial attempts to separate air with membranes were aimed primarily at the production of oxygen, and date back over a century (Prasad et al., 1994). However, the oxygen-selective

-240~

Separator

Heat Exchanger

Separator I I

I

I

I I

I

I

] Heat ] Exchanger

Feed 2.1% He 465 psia Residue Gas Residue Gas From PSA Units---~.~r~l[__ ~

To PSA 95% He 640 psia

Figure 15-22. Cryogenic/Membrane hybrid process for removing helium from natural gas. (Choe et al., 1988)

Membrane Permeation Processes

1283

nature of currently available polymeric membranes makes them more suitable for the production of nitrogen. In a typical operation, a single permeation stage can produce relatively pure nitrogen as residue gas at elevated pressure; while the permeate is oxygen-enriched air at atmospheric pressure. As a result of the process advantages of nitrogen production, and the rapidly developing commercial market for nitrogen, the development and use of membrane systems for nitrogen production have expanded rapidly.

Nitrogen In a paper prepared in 1993, it was estimated that the production of nitrogen from air had reached well over 2,000 tons per day, and more than 1,000 commercial units had been installed worldwide (Prasad et al., 1994). The practical economic limit for nitrogen purity has risen from 95-97 to 99.9%, with the ability to produce 99.9995% nitrogen from compressed air (Rice, 1990). However, extremely high purities require the use of a hybrid system employing a chemical deoxygenation step following the membrane unit. Nitrogen recoveries have gone from 22 to 30% (Anon, 1990). The major constituents in air, other than nitrogen and oxygen, are water vapor and carbon dioxide. Both gases are faster gases than oxygen and are concentrated in the oxygen permeate stream. Therefore, it is possible to obtain nitrogen with a dew point less than 40~ and a carbon dioxide concentration less than 20 ppm using membrane separators. Nitrogen is presently produced via cryogenic, PSA, and membrane systems. It is produced on site or delivered as liquid or as gas in cylinders. Figure 15-23 shows the typical economic areas of application of various nitrogen supply systems based on the quantity and purity of nitrogen required.

100 I

I

101 I IIIIIII

I

FLOW RATE, SIO FTS/H 102 103 104 10s I IIIII1(

I

I I!11111

I

IIIIIII

''i

I

I I

IIIII

!

I III11~

'

99

98 97 ~EUVER

95 MEMBRANE

94

Figure 15-23. Economic areas of applicability of nitrogen supply systems (Beaver et al., 1988). Reproduced with permission from AICHE Symposium Series, copyright 1988, American Institute of Chemica/ Engineers

1284

Gas Purification

The primary use for nitrogen produced by membrane separation is inert gas blanketing. Most flammable gases require a minimum of 10-12 vol% oxygen to sustain combustion. Since high purity nitrogen is not required for this application, 95-97% nitrogen from a membrane unit is satisfactory for inerting storage tanks. A membrane system to produce nitrogen for inerting the fuel tanks of military aircraft was described by Bhat and Beaver (1988) and Gollan and Kleper (1987). Depending on the type of aircraft and the flight conditions, bleed air from the jet engine compressors is available at 25 to 150 psig and temperatures up to 450~ The air is cooled to 70~176 depending on the membrane limitations, filtered to remove particulates, and then fed to the membrane unit. The inert gas produced has a nitrogen concentration of 88 to 97%. Crude and LNG tankers began inert blanketing of cargo tanks for safety in the early 1970s. Chemical tankers require a dry, non-reactive gas for purging lines and for preventing condensation inside cargo tanks. Metzger et al. (1984) reviewed the design of a system installed onboard two ships. The system was designed to produce a 95 to 99% nitrogen stream. Operating data for the system are presented in Table 15-17. Dry instrument gas can also be provided by attaching a connection to one or more of the membrane modules. Bhat and Beaver (1988) also reported on the use of membranes to control the composition of the atmosphere in the storage and transportation of fruit and vegetables. To slow respiration, the oxygen content of the atmosphere needs to be in the 1 to 5% range. However, two metabolic products, carbon dioxide and ethylene, also must be controlled to prevent accelerated ripening and decay. Membranes can provide primary control of the oxygen and secondary control of the carbon dioxide and ethylene. Although the resultant gas mixture may not be the optimum, the cost is less than that required to produce an exact optimum mixture of the gases. The use of a membrane system to replace delivered liquid nitrogen in a powdered metal sintering operation was described by McGinn and Pfitzinger (1988). The sintering furnace consisted of three zones: presintering, high heat, and cooling. A blend of 20% dissociated ammonia and 80% nitrogen from a liquid storage system was being used. In the new system,

Table 15-17 Operating Data for Ship Onboard Inerting System Case 1

Case 2

Input Flow, scfh Pressure, psig Temp,~

6674 441 122

4026 441 122

Inert Gas Flow, scfh Press, psig Temp,~ Dew point,~ % Nitrogen

3320 419 122 -85 95

989 419 122 -85 99

Source: Metzger et al. (1984)

Membrane Permeation Processes

1285

500 scfh of membrane-separated nitrogen was used in the presintering and cooling zones and dissociated ammonia was used in the high heat zone. The membrane-separated nitrogen contained a small amount of oxygen, which proved to be an advantage in the presintering zone by aiding in the removal of lubricants. The removal of the lubricants with the membraneseparated gas was far more effective than with the previously used nitrogen atmosphere. In addition to improved quality, a reduction in operating costs was observed. Delivered liquid nitrogen cost, including storage was $0.53 per hundred scf, while membrane gas production cost was only $0.28 per hundred scf. Beaver et al. (1986) reported the results of a feasibility study for generating 600 MMscfd of nitrogen on a ship moored to an offshore platform. The nitrogen generated was to be injected into an offshore oil reservoir to arrest the sinking of the seabed and the production platform fixed to the seabed. In the proposed process, turbine exhaust gases are cooled in a heat recovery boiler and an exhaust gas cooler. The gases are then preheated to prevent condensation and fed to a membrane unit. The nitrogen-rich residue stream from the membrane unit contains 2% oxygen and about 1% CO2. The oxygen is reduced to less than 5 ppm by catalytic reduction. A glycol dehydration system is used to remove water down to a - 4 ~ dew point. Product nitrogen is produced at 68~ and 900 psia. The system has three identical and independent process trains, which include nine gas turbines, nine heat-recovery units, six compressor systems, and six membrane-separation units. The heat recovery units produce about 600,000 pounds per hour of steam. It was concluded that the membrane process would be an energy efficient, low cost method of generating nitrogen. A comparison of the capital and operating costs of a membrane system versus PSA for producing nitrogen was developed by Gollan and Kleper (1986). The results are presented in Table 15-18. The analysis shows that the cost per standard cubic foot of nitrogen produced is the same for both systems, but the capital cost of the PSA system is one-third higher. The integration of membranes and catalytic oxygen removal to produce 99.9995% nitrogen was described by Beaver et al. (1988). Compressed air at 125 psig and 100~ is fed to a membrane separator that produces a 99.5% nitrogen stream. The nitrogen stream, at 110 psig, is then fed to a catalyst module where the nitrogen is mixed with a small amount of hydrogen and passed over a palladium catalyst. The catalytic reaction of oxygen and hydrogen to form water reduces the nitrogen stream oxygen content to less than 5 ppm. Beaver et al. (1988) also described a small scale application that integrates membranes and cryogenics to produce 5 liters per day of liquid nitrogen. A 100 psig compressed and filtered air stream is fed to a membrane unit that produces a 99% nitrogen stream. This stream is then sent to a cold head that is located inside a liquid nitrogen Dewar flask. The membrane module, made of hollow fibers, is two inches in diameter by three feet long. The unit can be reconfigured to produce liquid oxygen by changing the internal piping. High purity oxygen can be produced by adding a compressor and a second-stage membrane module. McReynolds (1985) described a commercially available package unit producing up to 99% nitrogen and up to 40% oxygen. The unit is sized to deliver up to 15,900 cfh of 95% nitrogen at 105 psi and 77~ Depending on the operational mode and capital depreciation option selected, the total operating cost ranges between $0.071 to $0.15 per hundred cubic feet, assuming $0.05/kWh electrical cost.

Oxygen Oxygen is presently produced via cryogenic distillation, PSA, vacuum swing adsorption (VSA), and membrane systems. It is produced on site or delivered as liquid or as gas in

1286

Gas Purification

Table 15-18 Operating Cost Comparisonfor 95% Nitrogen Generation Membranes Capacity, ton/day Capital Cost, FOB Installation, % Installed Cost, $ Operating Costs, S/day Membrane Replacement Power Cost Cooling Water Cost Operator Labor Cost Maintenance & Taxes Capital Charges Depreciation Total Operating Costs S/day S/ton $/100 SCF

3 $75,000 20 $90,000

PSA 3 $100,000 20 $120,000

16.01 35.10 0.43 4.30 3.89 32.93 12.96

0.00 40.50 0.43 4.30 5.18 43.90 17.28

105.61 35.20 0.13

111.59 37.20 0.13

Note: Membrane performance and costs based on A/C Technology Units. Source: Gollan and Kleper (1986)

cylinders. Figure 15-24 shows the market economics relative to size and purity. Note that the use of oxygen begins at 21%, its atmospheric concentration, while applications using nitrogen require concentrations beginning at 90%. A single-stage membrane yields oxygen concentrations up to 40%. Multi-stage systems can produce oxygen concentrations up to 60%; however, the economics are much less attractive. Oxygen passes through the membrane in the N2/O2 separation. Therefore, the feed gas (air) must be compressed, and the oxygen enriched air is produced at low pressure. Oxygen-enriched air produced by membranes has applications that include industrial combustion efficiency improvement, fermentation efficiency improvement, respiratory care, and pulp bleaching. Although these applications hold promise, commercial implementation has been slow. The major potential market is combustion enhancement, but progress has been limited by combustion system development. There is also limited use in respiratory care, particularly in Japan (Spillman, 1989). One advantage membranes have in medical applications is elimination of the humidification step. Oxygen produced via adsorption (PSA or VSA) or cryogenics is dry and requires humidification before patient use. Since water is a fast gas, it permeates with the oxygen in a membrane system. An application that integrates membranes with PSA is an extension of a previous nitrogen application--fuel tank blanketing on board military aircraft. The oxygen-rich permeate is fed to a PSA unit to produce 90% oxygen for crew breathing (Beaver et al., 1988). A comparison of oxygen-enriched air (OEA) produced by membranes and PSA was performed by Gollan and Kleper (1986). Their data are presented in Table 15-19. Membranes show a significant advantage in capital cost and power consumption.

Membrane Permeation Processes

1287

Table 15-19 Operating Cost Comparison for 35% Oxygen Enriched Air Generation PSA

Membranes

$250,000 15 $287,500

$480,000 15 $552,000

Operating Costs, S/day Membrane Replacement Power Cost Operator Labor Cost Maintenance & Taxes Capital Charges Depreciation

38.11 85.50 8.60 9.91 105.18 33.03

0.00 130.50 8.60 19.02 201.95 63.41

Total Operating Costs S/day S/ton

280.33 28.03

423.49 42.35

Capital Cost, FOB Installation, % Installed Cost

Source: Gollan and Kleper (1986)

FLOW - PURE OXYGEN EQUIVALENT, STD FT31H 100 101 102 103 104 105 106

I

i

80

e, e oe

70

[

:

! V,~

~

'

PSA

e

~ 5o

I O.~,~.E

e e o e e

60

o=~.

,. . . . . . . . . . . . . ----...~ .... I" ,' I | i

90

"

I

'i

I

e e e e o e |

30

MEMBRANE e e e e'

I

20

(Beaver et al., 1988). Reproduced with permission from AICHE Symposium Series, copyright 1988, American Institute of Chemical Engineers Figure 15-24. Economic areas of applicability of oxygen supply systems

1288

Gas Purification

SolventVapors Although permeability data on organic vapors and air were available in the early 1970s (Spangler, 1975; Rogers et al., 1972) process development did not begin until the mid-1980s. Operating conditions for membranes used in vapor recovery applications are quite different from the conditions seen in hydrogen or carbon dioxide applications. Inlet pressures are typically less than 125 psig, and a vacuum pump is frequently used on the permeate side to increase the pressure ratio. The permeate is often compressed for recovery of the solvents by condensation. Peinemann et al. (1986) described the use of membranes to recover solvent from an industrial oven. Hot solvent-laden air from the oven is fed to a blower to boost its pressure to 17.6 psia and then to a membrane system. A vacuum pump is used to maintain a low pressure on the permeate side of the membrane. The solvent-rich permeate is compressed and the solvent is condensed and recovered. The hot-solvent depleted residue gas is recycled back to the oven to reduce energy consumption. The operating and cost data for this system are presented in Table 15-20. In the manufacture of polyvinyl chloride (PVC), an offgas stream is produced that contains unreacted vinyl chloride monomer. This stream is usually compressed and condensed to recover as much monomer as possible. However, the vent stream from the condenser still contains a significant amount of monomer. Baker et al. (1991) described an installation that recovers 100 to 200 lb/hr of monomer from a condenser vent stream. The process vent stream is compressed to 65 psig and sent to a condenser operating at 14~ The condenser vent stream, containing 50% vinyl chloride monomer, is sent to the membrane unit. The

Table 15-20 Membrane Economics for Solvent Vapor Recovery Plant Operating Characteristics Concentration of Solvent in Feed Air Concentration of Solvent in Residue Concentration of Solvent in Permeate Volume Flow of Feed Air @ 17.6 psi Pressure Ratio

0.5 vol% 0.1 vol% 4.4 vol% 1400 scfm 0.05

Capital Cost of Plant Membranes Other Blower, 17 Hp Vacuum Pumps, 200 Hp Total

$40,500 $40,500 $30,000 $220,000 $331,000

Operating Costs (Annual Basis) Annual Fixed Costs Membrane Replacement Utilities ($0.05/kwh) Total

$48,000 $~3,5oo $71,ooo $132,5oo

Note: 1,000 L/day solvent recovery, 0.5 vol% solvent in feed. Source: Peinemann et al. (1986)

Membrane Permeation Processes

1289

monomer in the vent gas is reduced to less than 1%. The discharge of the vacuum pump on the permeate stream is recycled back to the inlet compressor suction. Baker et al. (1991) also described the installation of a membrane system on a CFC-11 and CFC-113 drum filling line. The unit is designed to process 10 scfm of saturated vent gas from the drum filling operation. The gas passes through a drier and then to a condenser operating at 5~ The condenser reduces the CFC concentration in the vent gas from 65 to 21%. A single-stage membrane is then used to reduce the CFC-11 concentration to 1.2%. Operating data and system costs are presented in Table 15-21, and the unit is pictured in Figure 15-25. The unit recovers approximately two pounds of CFC per 55 gallon drum filled.

Table 15-21 Vapor Recovery Membrane System Performance and Costs System Performance Condenser inlet solvent concentration Condenser temperature CFC removal: Condenser only Membrane separator + condenser

CFC-11

CFC-113

65% -15~

25% -15~

82% 98%

68% 96%

System Costs Capital cost Operating costs

$ 45,000 $ 8,500/yr $0.01/lb of CFC recovered

Source: Baker et al. (1991)

Figure 15-25. CFC vapor recovery system. Courtesyof Membrane Technologyand Research, Inc.

1290

Gas Purification

The use of membranes to recover HCFC-123 from the drying chamber of a film coating operation was described by Baker et al. (1992). A flow diagram of the system is shown in Figure 15-26. The vapor stream exiting the drying chamber is compressed to 125 psig, dried to prevent icing in the condenser, and cooled to 5~ where most of the HCFC-123 is condensed out. The condenser vent is routed to a membrane unit where the HCFC-123 concentration is reduced to less than 100 ppm. The unit recovers 15 pounds per hour and will pay for itself in 3,000 hours of operation based on $5/lb for HCFC-123. The system is pictured in Figure 15-27. Other solvent vapor applications for membranes include recovery of refrigerants from the purge stream of low temperature chiller refrigeration systems (Wijmans et al., 1991) and reducing emissions from industrial sterilizers (Baker and Wijmans, 1991), hydrocarbon storage tanks, dry cleaning operations, and printing and coating processes (Baker, 1985).

~.

Compressor

Feed HCFC-123 6.3% 50 scfm

f I

Condenser 1~ T

Residue

~

Ib HCFC-123 0.01% ~.9 ~m

33 kg/h HCFC-123

Permeate Vacuum Pump

Figure 15-26. Flow diagram of membrane system for the recovery of HCFC-123 vaporized in a film drying operation. (Baker et al., 1992)

Figure 15-27. HCFCvapor recovery system. Courtesyof Membrane Technologyand Research, Inc.

Membrane Permeation Processes

1291

REFERENCES Anon., 1990, "Membranes Shoot for the Big Time," Chem. Eng., April, pp. 37-43. Anon., 1985, "Amoco Starts Up CO2 Recovery Plant in Big West Texas Field," Sept. 9, p. 80. Babcock, R. E., Spillman, R. W., Godden, C. S., and Cooley, T. E., 1988, "Natural Gas Cleanup: A Comparison of Membrane and Amine Treatment Processes," Energy Prog., September, pp. 135-142. Baker, R. W., and Wijmans, J. G., 1991, "Process for Reducing Emissions from Industrial Sterilizers," U.S. Patent No. 5,069,686, December. Baker, R. W., Noriaki, Y., Mohr, J. M., and Kahn, A. J., 1986, "Separation of Organic Vapors from Air," Presented at the American Chemical Society Regional Meeting, Denver, CO, June 8-12. Baker, R. W., Kaschemekat, J., Wijmans, J. G., and Simmons, V. L., 1992, "Membrane Vapor Separation Systems for the Recovery of VOCs," Presented at the Air & Waste Management Association Annual Meeting, Kansas City, MO, June 21-26. Baker, R. W., 1985, "Process for Recovering Organic Vapors from Air," U.S. Patent No. 4,553,983, Nov 6. Baker, R. W., Kaschemekat, J., Simmons, V. L., and Wijams, J. G., 1991, "Membrane Pervaporation and Vapor Separation Systems for the Control of VOCs," Presented at the Ninth Annual Membrane Technology/Planning Conference, Nov. 6. Beaver, E. R., Graham, T. E., Johannessen, T., and Kvivik, H., 1986, "Inert Gas Generation Systems for Offshore Platforms," Energy Prog., September, pp. 149-154. Beaver, E. R., Bhat, P. V., and Sarcia, D. S., 1988, "Integration of Membranes with Other Air Separation Technologies," AIChE Symposium Series, No. 261, Vol. 84, pp. 113-123. Bhat, P. V., and Beaver, E. R., 1988, "Innovations in Nitrogen Inerting Using Membrane Systems," AIChE Symposium Series, No. 261, Vol. 84, pp. 124-129. Bhide, B. D., and Stern, S. A., 1993, "Membrane Processes for the Removal of Acid Gases From Natural Gas; I. Process Configurations and Optimization of Operating Conditions; II. Effects of Operating Conditions, Economic Parameters, and Membrane Properties," J. Memb. Sci., Vol. 81, Part I, pp. 209-237; Part II, pp. 239-252. Bollinger, W. A., Long, S. P., and Metzger, T. R., 1984, "Optimizing Hydrocracker Hydrogen," Chem. Eng. Prog., May, pp. 51-57. Bollinger, W. A., MacLean, D. L., and Narayan, R. S., 1982, "Separation Systems for Oil Refining and Production," Chem. Eng. Prog., October, pp. 27-32. Boustany, K., Narayan, R. S., Patton, C. J., and Stookey, D. J., 1983, "Economics of Removal of Carbon Dioxide from Hydrocarbon Gas Mixtures," Proceedings of the SixtySecond Gas Processors Association Annual Convention, pp. 146-149. Burmaster, B. M., and Carter, D. C., 1983, "Increased Methanol Production Using Membrane Separators," paper presented at the AIChE National Meeting, Houston, TX, March 27-31. Choe, J. S., Agrawal, R., Auvil, S. R., and White, T. R., 1988, "Membrane/Cryogenic Hybrid Systems for Helium Purification," Proceedings of the Sixty-Seventh Gas Processors Association Annual Convention, pp. 251-255. Coady, A.B., and Davis, J. A., 1982, "CO2 Recovery by Gas Permeation," Chem. Eng. Prog., October, pp. 44-49.

1292

Gas Purification

Cooley, T. E., and Dethloff, W. L., 1985, "Field Tests Show Membrane Processing Attractive," Chem. Eng. Prog., October, pp. 45-50. Cooley, T.E., 1990, "The Use of Membranes for Natural Gas Purification," Presented at the Gas Processors Association Meeting, European Chapter, Biarritz, France, May 17 and 18. Cutler, G., and Johnson, J., 1985, "Large-Scale CO2 Recovery with Membranes," Proceedings of the 1985 Laurence Reid Gas Conditioning Conference Proceedings, Univ. of Oklahoma, Norman, OK, pp. F1-F11. Cynara, 1995, "Gas Membrane CO2 Separations," Brochure: Experience List for CO2 Membrane Applications, Cynara Co., Houston, TX, Rev. 8/95 jah, Sept. 6. DaCosta, A. R., Fane, A. G., Fell, C. J. D., and Franken, A. C. M., 1991, "Optimal Channel Spacer Design for Ultrafiltration," J. Memb. Sci., Vol. 62, pp. 275-291. DaCosta, A. R., Fane, A. G., and Wiley, D. E., 1994, "Spacer Characterization and Pressure Drop Modelling in Spacer-Filled Channels for Ultrafiltration," J. Memb. Sci., Vol. 87, pp. 79-98. Donohue, M. D., Minhas, B. S., and Lee, S. Y., 1989, "Permeation Behavior of Carbon Dioxide-Methane Mixture in Cellulose Acetate Membranes," J. Memb. Sci., Vol. 42, p. 197. Doshi, K. J., Werner, R. G., and Mitariten, M. J., 1989, "Integration of Membrane and PSA Systems for the Purification of Hydrogen and Production of Oxo-Alcohol Syngas," AIChE Symposium Series No. 272, Vol. 85, pp. 62-67. Ettouney, H. M., A1-Enezi, G., and Hughes, R., 1995, Modeling of Enrichment of Natural Gas Wells by Membranes," Gas Sep. Purif, Vol. 9, No. 1, pp. 3-11. Flynn, A. J., 1983, "Wasson Denver Unit--CO2 Treatment," Proceedings of the Sixty Second Annual Convention of the Gas Processors Assoc., pp. 142-145. Fournie, F. J. C., and Agostini, J. P., 1987, "Permeation Membranes Can Efficiently Replace Conventional Gas Treatment Processes," J. of Pet. Tech., June, pp. 707-712. Funk, E. W., Kulkarni, S. S., and Swamikannu, A. X., 1986, "Effect of Impurities on Cellulose Acetate Membrane Performance," AIChE Symposium Series, No. 250, Vol. 82, pp. 27-34. Goddin, C. S., 1982, "Comparison of Processes for Treating Gases with High CO2 Content," Proceedings of the Sixty-First Annual Gas Processors Association Annual Convention, pp. 60-68. Gollan, A., and Kleper, M. H., 1986, "Membrane-Based Air Separation," AIChE Symposium Series, No. 250, Vol. 82, pp. 35-47. Gollan, A., and Kleper, M. H., 1987, "State-of-the-Art: Gas Separation," Proceedings of the Fifth Annual Membrane Technology~Planning Conference, Boston, MA, October, pp. 145-160. Graham, T., 1866, Philos. Mag., Vol. 32, p. 401. Hamaker, R. J., 1991, "Evolution of a Gas Separation Membrane 1983-1990," Presented at the International Conference on Effective Industrial Membrane Processes--Benefits and Opportunities, Edinburgh, Scotland, March 19-21. Heyd, J., 1986, "Hydrogen Recovery Using Membranes in Refining Applications," Presented at the National Petroleum Refiners Association Annual Meeting, Los Angeles, CA, March 23-25. Hogsett, J. E., and Mazur, W. H., 1983, Hydro. Process., August, p. 52.

Membrane Permeation Processes

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Houston, C. D., and Schell, W. J., 1986, "Recovery of Methane from Landfill Gas by the SEPAREX Membrane Process," Proceedings of Energy from Biomass and Wastes X, Institute of Gas Technology, Chicago, IL, April 7-10. Koros, W.J., and Fleming, G. K., 1993, "Review, Membrane-Based Gas Separation," J. Memb. Sci., Vol. 83, pp. 1-80. Koros, W. J., and Chern, R. T., 1987, "Separation of Gaseous Mixtures Using Polymer Membranes," Chapter 20 In Handbook of Separation Process Technology, R.W. Rouseau, ed., New York: John Wiley and Sons, pp. 862-953. Lacey, R. E., and Loeb, S., 1979, Industrial Processing with Membranes, New York: Robert E. Krieger Publishing Company, pp. 279-339. Lane, V. O., 1993, "Plant Permeation Experience," Hydro. Process., August, p. 56. Lee, A. L., Feldkirchner, H. L., Stern, S. A., Houde, A. Y., Gamez, J. P., and Meyer, H. S., 1995, "Field Tests of Membrane Modules for the Separation of Carbon Dioxide from Low-Quality Natural Gas," Gas Sep. Purif, Vol. 9, No. 1, pp. 35-45. Lee, S. Y., Minhas, B. S., and Donohue, M. D., 1988, "Effect of Gas Composition and Pressure on Permeation Through Cellulose Acetate Membranes," AIChE Symposium Series, No. 261, Vol. 84, pp. 93-101. Li, K., Acharya, D. R., and Hughes, R., 1990, "Membrane Gas Separation with Permeate Purging," Gas. Sep. Purif, No. 4, p. 81. Loeb, S., and Sourirajan, S., 1960, University of California, Los Angeles, Report No. 60-60. MacLean, D. L., and Narayan, R. S., 1982, "PRISM Separator Applications--Present and Future," In Hydrogen Energy Progress IV, Vol. 2. T. N. Verizoglu, W. D. Van Vorst, and J. H. Kelly, eds., New York: Pergamon Press, pp. 837-847. MacLean, D. L., Prince, C. E., and Chae, Y.C., 1988, "Energy--Saving Modifications in Ammonia Plants," Chem. Eng. Prog., March, pp. 98-104. MacLean, D. L., Stookey, D. J., and Metzger, T. R., 1983, "Fundamentals of Gas Permeation," Hydro. Process., August, pp. 47-51. MacLean, D. L., and Graham, T.E., 1980, "Hollow Fibers Recover Hydrogen," Chem. Eng., Feb. 25, pp. 54-55. Marquez, J. J., 1991, "Application of Hollow Fiber Membranes with High Permeate Pressures," Proceedings of the Seventieth Gas Processors Association Annual Convention, San Antonio, TX, March 11-12, pp. 212-221. Marquez, J. J., and Hamaker, R. J., 1986, "Development of Membrane Performance During SACROC Operations," Presented at the AIChE Spring National Meeting and Petro. Expo '86, New Orleans, LA, April 6-10. Mazur, W. H., and Chan, M. C., 1982, "Membranes for Natural Gas Sweetening and CO2 Enrichment," Chem. Eng. Prog., October, pp. 38-43. McCann, P., Ryan, J. M., and O'Brien, J. V., 1987, "The Mitchell Alvord South CO2 Plant Propane Recovery--Ryan Holmes," paper 74A presented at the AIChE 1987 Spring National Meeting, Houston, TX, March 29-April 2. McGinn, K. S., and Pfitzinger, M. S., 1988, "Producing a Sintering Atmosphere Through a Hollow Fiber Membrane," Proceedings of the International Powder Metallurgy Conference, Orlando, FL, pp. 229-241. McKee, R. L., Changela, M. K., and Reading, G. L., 1991, "CO2 Removal: Membrane Plus Amine," Hydro. Process., April, pp. 63-67.

1294

Gas Purification

McReynolds, K. B., 1985, "A New Air Separation System," Chem. Eng. Prog., June, pp. 27-29. MEDAL, (An Air Liquide group company), 1996, "Hydrogen (Medal)," Gas Processes '96, Hydro. Process., April, p. 119. Metzger, T. R., Handermann, A. C., and Stookey, D. J., 1984, "Shipboard Inert Gas Generation," Presented at the AIChE 1984 Summer National Meeting, Philadelphia, PA, Aug. 22. Meyer, H. S., and Gamez, J. P., 1995, "Gas Separation Membranes: Coming of Age for Carbon Dioxide Removal from Natural Gas," Proceedings of the 45th Annual Laurance Reid Gas Conditioning Conference, Norman, OK, Feb. 26-March 1, pp. 284-306. Mitchell, J. K., 1831, J. Roy. Inst., Vol. 2, No. 101, p. 307. Parro, D., 1984, "Membrane CO2 separation proves out at Sacroc tertiary recovery project," Oil & Gas J., Sept. 24, pp. 85-88. Peinemann, K. V., Mohr, J. M., and Baker, R. W., 1986, "The Separation of Organic Vapors from Air," AIChE Symposium Series, No. 250, Vol. 82, pp. 19-26. Poffenbarger, G. L., and Gastinne, P., 1989, "Hydrogen Membrane Applications and Design Considerations," Presented at the AIChE National Meeting, Houston, TX, April 2-6. Prasad, R., Notaro, F., and Thompson, D. R., 1994, "Evolution of Membranes in Commercial Air Separation," J. Memb. Sci., Vol. 94, pp. 225-248. Price, B. C., and Gregg, F. L., 1983, "CO2/EOR: from source to resource," Oil and Gas J., August 22, pp. 116-122. Rice, A. W., 1990, "Process for Capturing Nitrogen from Air Using Gas Separation Membranes: Deoxygenation," U.S. Patent No. 4,894,068. Rogers, C. E., Fels, M., and Li, N. N., 1972, "Separation by Permeation Through Polymeric Membranes," In Recent Developments in Separation Science, Vol. II. Cleveland: Chemical Rubber Company, pp. 107-155. Russell, F. G., 1984, "Applications of the DELSEP Membrane System," Chem. Eng. Prog., October, pp. 48-52. Ryan, J. M., and Schaffert, F. W., 1984, "CO2 Recovery by the Ryan/Holmes Process," Chem. Eng. Prog., October, pp. 53-56. Schaffert, F. W., Wood, N. V., and O'Brien, 1986, "The Seminole San Andres Unit CO2 Recovery PlantmMeeting the Challenges of EOR Gas Processing," Proceedings of the Sixty-Fifth Annual Convention of the Gas Processors Assoc., pp. 86-89. Schell, W. J., 1983, "Membrane Use/Technology Growing," Hydro. Process., August, pp. 43-46. Schell, W. J., and Houston, C. D., 1985, "Refinery Hydrogen Recovery with SEPAREX Membrane Systems," Presented at the National Petroleum Refiners Association Annual Meeting, San Antonio, TX, March 24-26. Schell, W. J., and Hoernschemeyer, D. L., 1982, "Principles of Gas Separation," Presented at the AIChE Symposium, Anaheim, CA, June 7-10. Schell, W. J., and Houston, C. D., 1982, "Spiral-Wound Permeators For Purification and Recovery," Chem. Eng. Prog., Oct., pp. 33-37. Schendel, R. L., 1995, personal communication (consultant), Sept. 6. Schendel, R. L., and Seymour, J., 1985, "Take Care In Picking Membranes to Combine with Other Processes for CO2 Removal," Oil & Gas J., Feb. 18, pp. 84-86.

Membrane Permeation Processes

1295

Schendel, R. L., 1984, "Using Membranes for the Separation of Acid Gases and Hydrocarbons," Chem. Eng. Prog., May, pp. 39-43. Schendel, R. L., Mariz, C. L., and Mak, J. Y., 1983, "Is Permeation Competitive?," Hydro. Process., August, pp. 58-62. Schendel, R. L., and Nolley, E., 1984, "Commercial Practice in Processing Gases Associated with CO2 Floods," paper presented at the World Oil and Gas Show and Conference, Dallas, TX, June 4-7. Shaver, K. G., Poffenbarger, G. L., and Grotewold, D. R., 1991, "Membranes Recover Hydrogen," Hydro. Process., June, pp. 77-80. Shirley, J., and Borzik, D., 1982, "Hollow Fiber Gas Separator Boosts NH3 Output by 50 tpd," Chem. Process., January, pp. 30-31. Spangler, G. E., 1975, "Analysis of Two Membrane Inlet Systems on Two Potential Trace Vapor Detectors," Amer. Lab., Vol. 7, No. 7, pp. 36-8, 40-5. Spillman, R. W., 1989, "Economics of Gas Separation Membranes," Chem. Eng. Prog., January, pp. 41-62. Spillman, R. W., Barrett, M. G., and Cooley, T. E., 1988, "Gas Membrane Process Optimization," Presented at the AIChE National Meeting, New Orleans, LA, March 9. Spillman, R. W., and Cooley, T. E., 1989, "Membrane Gas Treating," Proceedings of the Sixty-Eighth Gas Processors Association Annual Convention, San Antonio, TX, pp. 186-194.

Stern, S. A., 1986, "New Developments in Membrane Processes for Gas Separations," In Synthetic Membranes, M.B. Chenoweth, ed., New York: MMI Press. Stookey, D. J., Patton, C. J., and Malcom, G. L., 1986, "Membranes Separate Gases Selectively," Chem. Eng. Prog., November, pp. 36-40. Tomlinson, T. R., and Finn, A. J., 1990, "H 2 Recovery Processes Compared," Oil & Gas J., Jan. 15, pp. 35-39. Tonen Technology, and UBE Industries, 1990, "UBE Polyimide Membrane System for Hydrogen Recovery and Purification of Hydrofiner Offgas," Unpublished report provided by UBE Industries, March 16. U.S. Dept. of Energy, 1989, "Membrane Separation Processes for Liquid Hydrocarbons and Gases in the Petrochemical Industry," DOE Contract No. DE-AC01-87CE40762, Report DOE/CH 10093-57; DE89009464, September. Weller, S., and Steiner, W. A., 1950, "Engineering Aspects of Separation of Gases--Fractional Permeation Through Membranes," Chem. Eng. Prog., November, pp. 585-590. Wijmans, J. G., Kaschemekat, J., and Baker, R.W., 1991, "A Membrane Process for the Recovery of Volatile Organic Compounds from Process and Vent Streams," Presented at the Air & Waste Management Association 84th Annual Meeting and Exhibition, Vancouver, British Columbia, June 16-21. Wood, N. V., O'Brien, J. V., and Schaffert, F.W., 1986, "Seminole CO2 flood gas plant has successful early operations," Oil and Gas J., Nov. 17, pp. 50-54. Yamashiro, H., Hirajo, M., Schell, W. J., and Maitland, C.F., 1985, "Plant Uses Membrane Separation," Hydro. Process., February, pp. 87-89. Youn, K. C., Poe, W. A., Sattler, J. P., and Inlow, H. L., 1987, "CO2-recovery plant key to West Texas EOR operations," Oil and Gas J., Nov. 23, pp. 76-78.