t , ' . I

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THE DESIGN AND COSlS OF MEMBRANE ORGANIC VAPOR/AIR SEPARATION SYSTEMS

J.G. Wijmans, J. Kaschemekat and R.W. Baker Membrane Technology and Research, Inc., Menlo Park, CA

ABSTRACT

For the past eight years, Membrane Technology and Research, Inc. (MTR) has been developing a membrane process to separate organic vapors from effluent air streams. This process is now at the demonstration stage, and a number of pilot plants and small commercial systems are in operation. The process appears to be particularly suited to the recovery of organics from relatively concentrated streams in the range of 0.5-10% organic solvent. In this concentration range, the principal competitive technologies are condensation, carbon adsorption and incineration. However, these well-developed processes have a number of drawbacks. In this paper, we will first describe the principles of the membrane process. The factors determining the design of systems will be discussed together with four specific application designs. Finally, the economics of the process will be compared to competitive technologies.

The vapor separation process is shown in its simplest form in Figure 1. A contaminated air stream is introduced into an array of membrane modules. Organic vapors are preferentially drawn through the membrane by a vacuum pump. The organic vapor is condensed and removed as a liquid. The purified airstream is removed as the residue. Transport through the membranes is induced by maintaining the vapor pressure on the permeate side of the membrane, lower than the vapor pressure on the permeate side of the feed air stream. This pressure difference is most commonly achieved by means of a vacuum pump. However, the feed stream can also be compressed. Air and organic vapor permeate the membrane at rates determined by their relative permeabilities and the pressure difference across the membrane, Because the membrane is 10-100 times more permeable to organic vapor than air, a significant enrichment of organic vapor on the permeate side of the membrane is achieved. Depending on the system design, between 90-99% of the organic vapor is removed from the feed air stream and a permeate stream, enriched 5- to 50-fold in organic vapor, is

Solvent. depleled slr

R Solvent In alr

t

t Llquld

solvent

M T R W - l S

Figure 1. Flow .diagram .of a membrane vapor separation process.

I

I . . I .

The model used to predict gas permeation behavior is the solution-diffusion model. In this model, it is assumed that gas at the high-pressure side of the membrane dissolves in the membrane material and diffuses down a concentration gradient to the low-pressure side of the membrane, where the gas is desorbed. It is also assumed that the gas phases on either side of the membrane are in thermodynamic equilibrium with their respective polymeric interfaces, and that the interfacial sorption and desorption processes are rapid compared to the rate of diffusion through the membrane, Thus, the rate-limiting step is diffusion. For simple gases, Fick's law leads to the equation,

which can be further simplified to

where J is the membrane flux (cmS(STP)/cm%ec), D is the diffusion coefficient of the gas in the membrane (cm2/sec) and is a measure of the gas mobility, t is the membrane thickness, k is the Henry's law sorption coefficient (cmS(STP)/cmScmHg) linking the concentration of the gas in the membrane material to the pressure in the adjacent gas, and Ap (cmHg) is the pressure difference across the membrane. P is the permeability (cmS(STP)/cm2.sec.cmHg), equal to the product Dk, and is a measure of the rate at which a particular gas moves through a membrane.

A measure of the ability of a membrane to separate two gases, ( I ) and (2), is the ratio of their permeabilities, a, called the membrane selectivity,

Some membrane materials, particularly rubbery polymers, possess an intrinsically high selectivity for organic solvents over air, and can therefore be used in a membrane separation process targeted at separating organic vapors from air. The type of separations that can be obtained with these membranes is illustrated in Figure 2, which shows data for the separation of CCISF (CFC-I 1) from nitrogen. The factors that enter into membrane permselectivity have been discussed e1sewhere.l Over the past eight years, we have examined a number of polymer materials and measured their permeabilities to commonly used solvent vapors. Our experience has shown that a membrane selectivity of greater than 10 is required to achieve an economical separation. Some representative selectivities for the MTR-100 membrane for solvent vapors over nitrogen are shown in Table 1. For the most commonly used solvents, the selectivities obtained are above 20.

2

I . .

1 .

110

60

75

I I I ,

- CFC-11 -

Permeate concentratlon 50

(vol%)

20

25

- -

I 1 I

B

.

0 0 10 20 30 40

Feed concentratlon (vol%)

MTA-0896RP

Figure 2. Separation of CFC-I I (CClyF) from nitrogen with the MTR-100 membrane.

3

Table 1. Selectivities of the MTR-100 Membrane to Common Industrial Solvent Vapors, Measured at Ambient Temperature

Vapor Membrane Selectivity

Octane 90 - 100

I , 1,2-Trichloroethane

Isopentane

Methylene chloride

CFC- I I (CCISF)

1.1 , 1 -Trichloroethane

Isobutane

Tetroh yd raf uran

CFC- 1 13 (C&l,F4)

Acetone

CFC- 1 I4 (CZC1,Fd)

Hafon-1301 (CFSBr)

60

30 - 60

50

45

30 - 40

20 - 40

20 - 30

25

15 - 25

10

3

Other factors also determine the amount of enrichment of organic vapor obtained in the permeate stream in a membrane process. The first is the extent of removal of organic vapor from the feed. When a given volume of the feed stream enters the membrane, it immediately begins to lose organic vapor as the organic vapor preferentially permeates the membrane. Thus, the Concentration of organic vapor in the feed stream decreases as it passes through the module. Similarly, the concentration of organic in the permeate stream will be highest at the entrance of the membrane module where the feed concentration is highest and lowest at the exit of the module where the feed concentration is lowest. The resultant overall concentration of the permeate stream exiting from the membrane module is then an average of this varying concentration across the length of the module, As the fractional removal of organic vapor from the feed stream is increased, the average concentration of organic vapor in the permeate decreases. As a result, if greater than 90% removal is required, a multistep membrane process is used with potential recycling of the permeate stream.

4

A third factor affecting the performance of a membrane process is the pressure ratio, 4, across the membrane, defined as

(4) total feed pressure (p')

4 - total permeate pressure (pa) I

The pressure ratio is important because the partial pressure of the organic vapor on the permeate side of the membrane can never exceed the partial pressure on the feed side. Thus, even for an infinitely selective membrane, the concentration of organic vapor on the permeate side of the membrane can never be greater than times the concentration in the feed. This is an important result because organic vapor membrane separation units are generally operated with a feed stream pressure between 1 and 2 atm and a permeate stream under vacuum. These operating conditions are used because the permeate stream has a much smaller volume than the feed stream and it requires less energy to lower the pressure of this stream than to increase the feed pressure. As a practical matter, therefore, the pressure ratio across the membrane is usually maintained between 30 and 50. This limits the enrichment of organic vapor obtained in a single-stage membrane system to the same range when the fractional removal of organic vapor from the feed air is very small, in the range of 0-10%. In practical industrial systems, however, organic vapor removals of 90% from the feed are often required. In this case, the enrichment of organic vapor obtained in a single-stage membrane system is in the range of 5-10. If higher fractional removal of solvent from the feed is required, combination membrane stages are used.

The membranes used in this vapor separation process are composites having the multilayer structure illustrated in Figure 3. The membrane consists of a tough, nonwoven polyester paper that provides mechanical strength. A finely microporous polymer layer then provides a smoother surface on which an ultrzthin, dense polymer is coated. I! is this ultrathin, dense layer that is responsible for the separation properties because the dense layer is very thin, of the order of 1 pm. These membranes have high fluxes. They are also much more permeable to organic vapors than to air.

Variable permsetectfve layer

- 5? & support Mlcroporous layer

@- Nonwoven polyester paper

MTR-062-F

Figure 3. Schematic of an MTR composite membrane. MTR produces these membranes in rolls 40 inches wide and 100-200 yards long.

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1 . . .

These membranes are fabricated into spiral-wound membrane modules. Industrial-sized modules are typically used in systems where the feedstream is a t atmospheric pressure or just a little above, and a pressure difference across the membrane is provided by drawing a partial vacuum on the permeate side. In these modules, the membrane is supported on the feed and permeate sides by plastic mesh spacers. The membrane and spacers are wound spirally around a central collection pipe. A schematic of a membrane module is given in Figure 4.

\ ,Spaerr

Figure 4. Schematic of a spiral-wound membrane module. Industrial modules contain between 4 and 6 m2 of membrane.

Svstem Desian

A number of system designs car. be used to obtain varying degrees of separation. In this illustration, each system design has been sized to treat a 100-scfm gas stream containing I vol% of CFC-I1 (CCI,F), which has a selectivity, a, over nitrogen of approximately 45. In the calculations illustrated in Figure 5 , the feed gas is compressed to 15 psig (150 cmHg abs) and a vacuum pump is used to generate a permeate pressure of 5 cmHg abs. The pressure ratio, 4, across the module is then 30; the performance of the membrane system is calculated using a computer model based on the computational method of Shindo.'

The single-stage system shown in Figure Sa is the simplest design. This system uses 48 m2 of membrane to achieve 90% removal of CFC-11 from the feed gas, producing a permeate containing 7.2% CFC-11. Because CFC- 1 1 is relatively volatile, cooling the permeate stream to 5°C only condenses a portion of the solvent and a substantial fraction is therefore recycled back to the feed. This recycle stream would become larger as the feed stream becomes more dilute. For this reason, single-stage systems are usually only used with streams in which the solvent concentration is already close to saturation.

Four system designs are illustrated in Figure 5.

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The size of the recycle stream can be reduced by using the two-stage system shown in Figure 5b. In this system, the permeate from the first stage becomes the feed to a second stage. This configuration allows an additional 6-fold enrichment of the organic solvent. Because the feedstream to the second stage is greatly reduced in volume, the size and cost of the second stage is normally only 10-20% as large as the first stage. It is possible to treat relatively dilute streams with two-stage systems and still obtain 90% vapor removal from the feed stream. For this reason, this design is preferred for many applications where the feed stream is relatively dilute and an enrichment of 20- to 50-fold is required to make condensation practical.

In some applications, more than 90% removal of organic vapor removal from the feed stream is required. In these applications, a multistep system is usually used. In this system, the first step removes 90% of the organic vapor from the feedstream, producing a permeate enriched 7-fold in CFC-11. The residue from the first step then becomes the feed to the second step where an additional portion of the CFC-I1 vapor is removed, achieving an overall vapor removal of 99%. The second step uses almost as much membrane area to remove the last 9% of the solvent vapor as the first step uses to remove the first 90%. The second membrane step, therefore, increases the total cost of the system by 1.5- to 2-fold. Because the permeate from the second step is relatively dilute, it is recycled to the feed of the first step.

A two-step design is illustrated in Figure 5c.

The final system design, illustrated in Figure 5d, is a two-step, two-stage system. This system achieves 99% removal of solvent vapor from the feed gas and achieves a substantial enrichment to the permeate in the second stage. As shown, the permeate from the first step is treated in a second stage to produce a concentrated stream which can be easily condensed to recover the organic vapor. The dilute permeate from the second step is recycled to the incoming feed stream.

7

0.1% CFC-11 87.4 rctm

~embronr unll 40 ma

100 u t m

1 7.2% CFGl I

12.6 Bctm

8 ) One-Stage System

0.1% CFC-11 90 rctm

Flrrl rtap. mombrano unll

53 m2 1% C f G I l loo whn

1% CFC-ii 12 rctm

0.01% CFC-11 DS.9 8 C h n

mombnnr unn

in CFGl l io0 rctm

7.0% CFC-11 14 rctm

Suond at- "bran. unll CFC-11

4 s d . wtm

4sn CFC-11 2 Utm

b) Two-Stage System

c) Two-step System

k c o n d rlrg. nwmbrrnr unll

4.7 rn2

0.01% CFC-11 OB rcfm

d) Two-Stage, Two-step System

Figure 5. Four membrane system configurations are illustrated for a 100-scfm air stream containing 1% CFC- 1 1 . (Absolute pressures: vacuum pumps, 5 cmHg; compressors, I50 cmHg.)

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The choice of system design depends on the concentration of the feed stream, the degree of separation required and the permselectivity and condensability of the organic vapor, In the examples shown in Figure 5, the feed and permeate pressures were fixed. In practice these would be vanied to obtain the optimum system design. Increased feed pressure generally leads to lower capital costs because the membrane area is reduced. However, the operating costs of high-pressure systems are higher because of the higher energy consumption, and thus, a balance must be struck between operating and capital cost. In the following sections, specific system designs for a number of applications are used to illustrate the potential of the process.

To illustrate the new membrane separation system's potential for waste gas utilization and reduction and energy savings, we use four practical examples, presenting an estimate of the costs of the process in each application. The assumptions used in the cost and system efficiency calculations are listed in Table 2,

Table 2. Parameters Used in Performance and Cost Calculations

Yearly operation 7,200 hrlyr

Module lifetime 3 years

Operation and maintenance

Electricity cost SO.OS/kWh

10% of capital costs

Compressor efficiency 60%

Vacuum pump efficiency 35%

Chiller performance Manufac tursr's specifications

Module cost $300-500/m' (volume- dependent)

Component cost Manufacturer's OEM price

Membrane flux and selectivity Laboratory module performance data

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The module lifetime of three years used in these calculations is based on an extrapolation of field data, Vapor separation with membranes is a new technology and systems have not yet been in operation for three years. However, other membrane gas separation systems have demonstrated membrane lifetimes of more than 5 years. No sign of our membrane's deterioration of performance has been observed in field demonstrations.

Vent Gas from Tank F i b ODeratim . . The first application where membrane vapor separation has advantages over

alternative technologies involves retrofitting an existing vent condenser with a membrane unit to improve the efficiency of the system. In this example, we consider a stream produced during CFC-I1 tank filling operations. The vent gas saturated with the CFC vapor is currently passed through a molecular sieve dehydrator and then to a -15'C condenser. The condenser reduces the CFC concentration substantially to approximately 21% CFC-11. In the past, this 21% stream would have been vented without further treatment. Now, however, recovery of the CFC in the vent is both economically and environmentally driven. A single-stage membrane unit easily reduces the vent gas CFC concentration by a further 90%, producing a CFC-enriched permeate that is recycled to the front of the condenser, The final CFC-I1 concentration in the vent gas is reduced from 21% to 1.2%. The flow chart and design calculations for this system are shown in Figure 6. This is a rather small system so the cost per square meter of membrane area is high. However, the cost of operating the unit is a t iny fraction of the value of the recovered CFC. The system pays for itself after only a few hundred hours of operation.

An alternative to the membrane process would be lowering the temperature of the condenser. However, to achieve 1.2% CFC-II in the final vent gas would require a condenser operating at -5O'C, at a cost substantially more expensive than the membrane unit.

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1.2% CFC-11 3.9 rctm

t Condenser

:15'c. 21% CFC-11 5.6 scfm 65% CFC-11

1 0 rctm r

I 1 I

Llquld CFC-11

I

62% CFC-11 1.7 scfm

Membrane selectivity 45 Membrane area 10 ma Feed pressure 80 cmHg abs Permeate pressure S cmHg abs

Capital Cost S 45,000

Annual Operating Cost Labor and maintenance S 4,SOO Module replacement 2,000 Energy 2,000

Total S 8,50O/yr

Operating Cost $0.01/lb of CFC-I 1 recovered

Figure 6. A membrane system for solvent recovery from tank filling operations. The system recovers 98% of the CFC-I1 contained in the original feed stream.

Methylene C h l o r i d e v a r v from a Pharmaceutical Co-

Methylene chloride is a commonly used solvent, especially in the pharmaceutical industry. Because methylene chloride is so volatile, the air streams produced are often quite concentrated. In many locations, these emissions are not regulated and the streams are vented without treatment. Conventional treatment processes for this stream are expensive. Incineration requires high-cost scrubbers to treat the chloride acid gases produced, carbon adsorption is marginally effective, and condensation is only effective if the gas is cooled to -4O'C and below. A membrane unit designed to recover 90% of a methylene chloride off-gas is illustrated in Figure 7. The capital cost of this system is S180,000, which is extremely competitive with alternative technologies. A n attractive feature of the system shown in Figure 7 is that the condenser operates at relatively modest temperatures of S'C. Not only does this lower the capital and operating costs of the system, but operational reliability is improved since periodic de-icing of the condenser is not necessary. A compression- condenser system designed to achieve this same removal would require cooling of the gas stream to -20' to -4O'C. This low temperature requires a two-stage chiller with concomitant excessive capital and operating costs.

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0.2% C*CI2 196 rchn

h Iembranr b

Vacuum Pump . 8 Compressor

2% CHS12

T' 10.4% CH#2

Condrnrer 5'C

Llquld CHsl2 52 lblh

Permeate from Epled - Residue

I96 * Fiow (scfm) 200 I

Concentration (%) 2 49.9 0.2

Membrane selectivity 30 Membrane area 109 Feed pressure 30 psig Permeate pressure 10 cmHg

Capital Cost S 185,000

Annual Operating Cost Labor and Maintenance S 18,500 Module replacement 18,300 Energy 12,200

Total S 49,00O/yr

Operating Cost S0.13/lb of methylene chloride recovered

Figure 7. A two-stage membrane system used to recover methylene chloride from a 2OO-scfm, 2% feed stream.

The system shown in Figure 7 is designed to achieve 90% removal of methylene chloride from a 200-scfm stream containing 2% solvent. In industrial processes, the solvent concentration will vary over a wide range. Here concentration changes will affect the ability of membrane and condensation systems to remove the solvent, The effect of these changes in feed gas concentration on system recovery are illustrated for both systems in Figure 8.

4 -

2 -

1 .

M.lhy(ono shlorldo tomoval

px) to .WC

I I I I I I 1 I 0

Lm4maW

Figure 8. Effect of changes in feed air vapor concentration on overall recovery for the processes illustrated in Figure 6.

Both the membrane and compression/condensation systems are less efficient at lower feed concentrations, However, even with compression to 150 psig and cooling to -2O'C, the compression-condensation system achieves no solvent removal once the feed concentration falls below 0.4%. At this lower concentration, the membrane sysSem stilt achieves 85% solvent iemova! and only falls below 50% removal when the feed concentration is reduced to 0.15% solvent.

A third application of a membrane vapor separation process is illustrated in Figure 9. In this example, the membrane unit is designed to fractionate a relatively concentrated CFC-I 1 stream produced as a process byproduct. This stream is too dilute for economical compression-condensation, but too concentrated to be sent directly to a carbon adsorber. Therefore, the stream is currently mixed with large volumes of more dilute streams and sent to a carbon adsorber. Even though this is a relatively small stream, it represents a significant fraction of the total mass sent to the carbon adsorber. A two-stage membrane system, shown in Figure 9, will split this stream into two fractions. The concentrated 99% fraction is now able to be compressed and condensed at moderate temperatures and pressures. The dilute 2% fraction places much less demand on a carbon adsorption system. The net result of isolating this stream is that a much smaller carbon adsorption system is required and regeneration cycles are lengthened. This approach matches the most appropriate recovery technology with those streams that can be isolated at their process sources. The capital and operating costs of the system compare favorably to existing technologies.

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. .

2% CFC-11

t .c 82m2 49 sctm 50% CFC-11

loo scfm , I

Permeate from ~~ Residue

Flow (scfm) 100 50 Concentration (%) 50 99

Membrane selectivity 45 Membrane area 94 ma Feed pressure 80 cmHg abs Permeate pressure 5 cmHg obs

Capital Cost $ 180,000

Annual Operating Cost Labor and Maintenance S 18,000 Module replacement 16,000 Energy 13,300

9 2

~

Total $ 47,30O/yr

Figure 9. A two-stage membrane system for fractionation of vapor streams for further processing.

Pecoverv of a CFC - 113 Off-Gas

CFC-113 is an important cleaning solvent material in the electronics industry. It is extremely volatile and is therefore difficult and expensive to recover by condensation or carbon adsorption. A two-step, two-stage system designed to treat a 50-scfm 3% CFC-113 stream is illustrated in Figure 10. Because of the high value of the solvent and its environmental impact, 95-99% CFC recovery is required. This means a more costly two-step system must be used. The second stage is required to concentrate the gas sufficiently to allow condensation at O'C. However, the second stage does not increase the cost of the unit substantially. A photograph of a two-step, two-stage system able to treat this type of stream is shown in Figure 11.

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0.07SW CFC-113 48.5 scfm

Compressor 8

3% CFC-113 50 rcfm

Llquld CFC-113 35 lblh

MTCI*OIS.r)

Flow (scfm) 50 Liquid 48.5 Concentration (%) 3.0 -- 0.075

Membrane selectivity 20 Membrane area 47 mz Vacuum pumps IS hp Compressors 9 hP

Capital Cost $ 150,000

Annual Operating Cost Labor and Maintenance $ 15,000 Module replacement 7,800 Energy 6,400

Total _ _ -

$ 29,20O/yr

Operating Cost $1.35/1,000 scf feed $0. I 1 /I b CFC- 1 1 3 recovered

Schematic design and estimated costs of a two-step system designed to recover 95% of CFC-I 13 from a 3% feedstream. The value of the recovered solvent far exceeds the operating cost per pound.

Figure 10.

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The economics of these procedures are appealing. Installation of this processes makes good environmental and economic sense. At today's tax levels, CFC-I 13 costs approximately S3.00/lb. Even if the value for the reclaimed solvent is ignored, the credit for the avoided tax produces a payback after 5,000 hours of operation.

8

b

Figure 11. Photograph of a two-stage, two-step membrane unit capable of treating 50-100 scfm of air containing solvent vapors.

Membrane vapor recovery is a new technology that is able to separate and recover organic solvent vapors from air. The process is particularly suited to the treatment of streams containing more than 0.5% solvent. With these streams, the value of the recovered solvent frequently exceeds the cost of the processes. A number of small plants and pitot units have been installed (August 1990). No major problems have emerged and the process appears to be reliable and efficient.

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References

1.

2.

3.

4.

3.

6.

7.

R.W. Baker, N. Yoshioka, J.M. Mohr and A.J. Khan, "Separation of Organic Vapors from Air," ' , 259 (1987).

K.V. Peinemann, J.M. Mohr and R.W. Baker, "The Separation of Organic

2201 (1970).

I

Liquids through Highly Swollen Polymer Membranes," J.. Polvm. Sc i. 14,

J.G. Wijmans and V.D. Helm, "A Membrane System for the Separation and Recovery of Organic Vapors from Gas Streams," AlChE SvnUZQsium Series fro. 272, Vol. 85, 74 (1989).

R.W. Baker, "Process for Recovering Organic Vapors from Air," U.S. Patent 4,353,983 (November 19, 1985).

J.G. Wijmans and B. Ahlers, "Membrane Process for Treatment of Fluorinated Hydrocarbon-Laden Streams," U.S. Patent 4,906,256 (March, 1990).

R.D. Behling, L. Ohlorogge, K.V. Peinemann and E. Kyburz, "The Separation of Hydrocarbons from Waste Vapor Streams," AIC hE SvmDos ium Ser ies No, U, Vol. 88, 68 (1989).

Y. Shindo, T. Habuta, H,. Yoshitome and H. Inous, "Calculation Methods for

459 (1985). Multicomponent Gas Separation by Permeation,. &R, Sci. Tec tmdLU , 445-

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BIOGRAPHIES

Hans Wijmaru received his Ph.D. in the membrane research group of Professor C.A. Smolders at the University of Twente, The Netherlands. He received the 1984 Royal Dutch/Shell Research Award for young scientists. Dr. Wijmans joined Membrane Technology and Research, Inc. in Menlo Park, California in 1984. He currently holds the position of Director of Research and oversees the company's activities in membrane-based gas and vapor separations and in pervaporation.

Jiirgen Kaschemekat received his Master's Degree in Chemical Engineering i n 1971 from the College of Technology in Aachen, West Germany. He then joined the membrane research group at the GKSS Research Center in Geesthacht, West Germany, where he worked for twelve years on the preparation of reverse osmosis membranes and systems for seawater desalination. He also designed and constructed pervaporation units for a variety of applications, including dealcoholization, production of ethanol from fermentation broths and treatment of wastewater. Mr. Kaschemekat came to MTR in 1987, and directs the company's system design group for vapor separations and pervaporation.

Richard Baker has specialized in various applications of membranes and membrane research since receiving his Ph.D. in Physical Chemistry in 1966 under one of the pioneers of membrane science, Professor R.M. Barrer, at Imperial College, London. Between 1966 and 1968 at Amicon Corporation, Lexington, MA, he developed a series of ultrafiltration membranes now sold under the name DiaflowQ. Later (1971-74), at Alza Corporation, Palo Alto, CA, he collaborated in the development of the Ocusert" ocular delivery system. He was co-founder and Director of Research at Bend Research, Bend, OR until founding Membrane Technology and Research, Inc. (MTR) in Menlo Park, CA in 1982. MTR was set up as a contract research company primarily specializing in membrane technology applications and now offers commercial membrane systems for recovering organic vapors from air and water streams. Dr. Baker is President of MTR.