Hollow fiber membrane contactors

Journal of Membrane Science 159 (1999) 61±106

Hollow ®ber membrane contactors Alan Gabelmana, Sun-Tak Hwangb,* b

a Givaudan Roure Flavors, Cincinnati, OH 45216, USA Department of Chemical Engineering, University of Cincinnati, Cincinnati, OH 45221-0171, USA

Received 26 March 1998; received in revised form 13 January 1999; accepted 13 January 1999

Abstract A membrane contactor is a device that achieves gas/liquid or liquid/liquid mass transfer without dispersion of one phase within another. This is accomplished by passing the ¯uids on opposite sides of a microporous membrane. By careful control of the pressure difference between the ¯uids, one of the ¯uids is immobilized in the pores of the membrane so that the ¯uid/¯uid interface is located at the mouth of each pore. This approach offers a number of important advantages over conventional dispersed phase contactors, including absence of emulsions, no ¯ooding at high ¯ow rates, no unloading at low ¯ow rates, no density difference between ¯uids required, and surprisingly high interfacial area. Indeed, membrane contactors typically offer 30 times more area than what is achievable in gas absorbers and 500 times what is obtainable in liquid/liquid extraction columns, leading to remarkably low HTU values. Although a number of membrane module geometries are possible, hollow ®ber modules have received the most attention. In general, tube side mass transfer coef®cients can be predicted with reasonable accuracy; on the other hand, shell side coef®cients are more dif®cult to determine, and several research groups are currently addressing this problem. Membrane contactor technology has been demonstrated in a range of liquid/liquid and gas/liquid applications in fermentation, pharmaceuticals, wastewater treatment, chiral separations, semiconductor manufacturing, carbonation of beverages, metal ion extraction, protein extraction, VOC removal from waste gas, and osmotic distillation. This paper provides a general review of hollow ®ber membrane contactors, including operating principles, relevant mathematics, and applications. # 1999 Elsevier Science B.V. All rights reserved. Keywords: Microporous and porous membranes; Organic separations; Membrane contactors

1. Introduction Gas/liquid and liquid/liquid contacting operations are traditionally done using some type of tower, column or mixer±settler. Usually, the main challenge in designing and operating these devices is to maximize the mass transfer rate by producing as much *Corresponding author. Tel.: +1-513-556-2691; fax: +1-513556-3473; e-mail: [email protected]

interfacial area as possible. For packed columns this requires judicious selection of packing material and uniform distribution of ¯uids before they enter the packed bed. Alternatively, for devices with moving internals the design challenge is to minimize the bubble or droplet size of the dispersed phase and maximize the number of bubbles or droplets. Although columns and other traditional ¯uid/¯uid contactors have been workhorses of the chemical industry for decades, an important disadvantage is

0376-7388/99/$ ± see front matter # 1999 Elsevier Science B.V. All rights reserved. PII: S 0 3 7 6 - 7 3 8 8 ( 9 9 ) 0 0 0 4 0 - X

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A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

the interdependence of the two ¯uid phases to be contacted, which sometimes leads to dif®culties such as emulsions, foaming, unloading and ¯ooding. An alternative technology that overcomes these disadvantages and also offers substantially more interfacial area than conventional approaches is non-dispersive contact via a microporous membrane. Using a suitable membrane con®guration such as a hollow ®ber or a ¯at sheet, ¯uids to be contacted ¯ow on opposite sides of the membrane and the ¯uid/¯uid interface forms at the mouth of each membrane pore. Mass transfer occurs by diffusion across the interface just as in traditional contacting equipment. However, unlike most membrane operations, the membrane imparts no selectivity to the separation, i.e., it has no impact on the partition coef®cients [28,61,98,109]. Furthermore, as opposed to more conventional membrane applications such as micro®ltration, ultra®ltration and reverse osmosis, the driving force for separation is a concentration rather than a pressure gradient; indeed only a very small pressure drop across the membrane is required to ensure that the ¯uid/¯uid interface remains immobilized at the mouth of the pore. Advantages membrane contactors offer over columns and other conventional mass transfer equipment include the following:  The available surface area remains undisturbed at high and low flow rates because the two fluid flows are independent. This is useful in applications where the required solvent/feed ratio is very high or very low. In contrast, columns are subject to flooding at high flow rates and unloading at low ones.  Emulsion formation does not occur, again because there is no fluid/fluid dispersion.  Unlike traditional contactors, no density difference is required between fluids; membrane contactors can accommodate fluids of identical density and can be operated in any orientation.  Scale-up is more straightforward with membrane contactors. That is, membrane operations usually scale linearly, so that a predictable increase in capacity is achieved simply by adding membrane modules (subject of course to the limitations of support equipment such as transfer pumps, piping, etc.). On the other hand, scale-up with conventional equipment is not nearly as straightforward.



Modular design also allows a membrane plant to operate over a wide range of capacities. That is, small or large capacities can be obtained simply by using few or many modules.  Aseptic operation is possible, a feature which can be advantageous in processes such as fermentation (see Section 6.5). For example, during fermentation the broth can be circulated through a membrane contactor then back into the fermenter, while a suitable extraction solvent passes through the contactor on the opposite side of the membrane. In this manner the product of interest is continuously removed, so that product inhibition is relieved and productivity is improved. In principle, a traditional contactor could be used here instead of a membrane contactor, but in practice aseptic operation is much more difficult with the former.  Membrane contactors can also be used to increase conversion with equilibrium-limited chemical reactions in general. That is, by circulating the reactor contents through the contactor against a suitable extraction solvent or stripping gas, the products are removed and the reaction equilibrium shifts to the right.  Interfacial area is known and is constant, which allows performance to be predicted more easily than with conventional dispersed phase contactors. On the other hand, interfacial area is quite difficult to determine in dispersive contactors because the bubble or droplet size distribution depends on operating conditions and fluid properties, which is why the mass transfer coefficient and interfacial area are usually lumped together (e.g., kGa, kLa) in mass transfer calculations. With packed columns the interfacial area per unit volume may be known, but it is often difficult to determine the loading, i.e., what fraction of the available surface is actually used.  Substantially higher efficiency (as measured by the height of a transfer unit or HTU) is achieved with membrane contactors than with dispersive contactors, as discussed in more detail below.  Solvent holdup is low, an attractive feature when using expensive solvents.  Unlike mechanically agitated dispersed phase columns, membrane contactors have no moving parts. On the downside, membranes also have disadvantages:

A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106



The membrane introduces another resistance to mass transfer not found in conventional operations: the resistance of the membrane itself. However, this resistance is not always important, and steps can be taken to minimize it.  Membrane contactors are subject to shell side bypassing, which results in a loss in efficiency. Often bypassing is not a problem in the laboratory but becomes an issue upon scale-up to larger contactors. Fortunately, several design improvements have been proposed to address this problem, as explained below (see Sections 5.2.3 and 5.2.4).  Membranes are subject to fouling, although this tends to be more of a problem with pressure-driven devices than with concentration-driven ones such as membrane contactors.  Membranes have a finite life, so that the cost of periodic membrane replacement needs to be considered.  The potting adhesive (e.g., epoxy) used to secure the fiber bundle to the tubesheet may be vulnerable to attack by organic solvents.  The achievable number of equilibrium stages is limited by pressure drop constraints, as explained below under Sections 3.2 and 5.4. These relatively few disadvantages are often outweighed by the numerous advantages cited above. For this reason, membrane contactors have attracted the attention of many interested parties from both academia and industry for a diverse range of applications. As a result, our level of understanding of the technology has improved remarkably since early studies of blood oxygenation were performed in the 1970s and early 1980s [38,135]. The discussion below provides a general review of membrane contactors with the primary focus on hollow ®ber modules since that con®guration has been used for the majority of work reported to date. Some applications of hollow ®ber membrane contactors have been reviewed by others and are only brie¯y mentioned here. Membrane distillation, a thermally driven process for transferring a solvent across a non-wetted membrane, was reviewed by Sirkar [123]. Membrane perstraction, a process where one or more components of a liquid feed solution diffuse across a non-porous membrane into a stripping or sweep liquid, was discussed by Sirkar [124]. Osmotic distillation is discussed brie¯y in Section 6.11.

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2. Commercially available modules for industrial applications Although mass transfer operations can be conducted using a number of different membrane con®gurations, including ¯at sheet, spiral wound, rotating annular, and hollow ®ber, the latter has received the most attention. Hollow ®ber modules targeted for industrial applications (as opposed to medical ones, e.g., blood oxygenation) are available from a variety of sources, although some are designed for pressuredriven ®ltration processes (see below) rather than concentration-driven mass transfer. The most well-known module designed for concentration-driven mass transfer is the Liqui-Cel1 Extra-Flow module offered by CELGARD LLC (Charlotte, NC; formerly Hoechst Celanese), shown in Fig. 1. This module uses Celgard1 microporous polypropylene ®bers that are woven into a fabric and wrapped around a central tube feeder that supplies the shell side ¯uid. Woven fabric allows more uniform ®ber spacing, which in turn leads to higher mass transfer coef®cients than those obtained with individual ®bers (see Section 5.2.4). Typically the ®ber inside diameter and wall thickness are 240 and 30 mm, respectively. The ®bers are potted into a solvent-resistant epoxy or polyethylene tubesheet, and the shell casing is polypropylene, PVDF or 316L stainless steel [116,117,126,138]. As shown in Fig. 1, the Extra-Flow module contains a central shell side baf¯e, a feature which offers two advantages. First, the baf¯e improves ef®ciency by minimizing shell side bypassing; second, it provides a component of velocity normal to the membrane surface, which results in a higher mass transfer coef®cient than that achieved with strictly parallel ¯ow (see Section 5.2). The smallest Liqui-Cel modules are 212 inches in diameter and contain 1.4 m2 of contact area, while the largest are 10 in. in diameter and offer 130 m2 of contact area by virtue of 225 000 ®bers. The large modules can handle liquid ¯ow rates of several thousand liters per minute [116,117,126,138]. Membrane Corporation (Minneapolis, MN) offers modules designed for bubble-free gas/liquid mass transfer, which provides advantages over conventional aeration methods in applications such as bioremediation and wastewater treatment (see Section 6.8). These modules ®t within standard PVC pipes and

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A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

Fig. 1. The Liqui-CelTM Extra-Flow membrane contactor from CELGARD LLC. The central baffle minimizes shell side bypassing and provides a component of velocity normal to the fibers (redrawn from [125]).

contain multiple ®ber bundles, each containing about 500 ®bers. The ®bers are potted into polyurethane at one end only, and individually sealed at the other end, so that there is no exhaust gas stream. That is, all entering gas exits by diffusing across the membrane into the surrounding water, leading to 100% gas transfer ef®ciency. The ®ber bundles are ¯uidized by the ¯owing liquid, and rapid ¯uidization is facilitated by the use of bundles which are ¯at instead of cylindrical. The packing density is only 10%; this provides plenty of room for ¯uidization, which in turn leads to high gas/liquid mass transfer rates, low pressure drop, and free passage of solids. The ®bers are composed of a 1 mm layer of polyurethane sandwiched between two layers of microporous polyethylene; inside and outside diameters are 220 and 270 mm, respectively [110]. W.L. Gore & Associates (Elkton, MD) also markets a module designed for bubble-free gas/liquid mass transfer primarily for ozonation of semiconductor cleaning water (see Section 6.10). Their DISSO3LVETM module features expanded polytetra¯ouroethylene (PTFE) ®bers that are compatible with highly corrosive chemicals such as ozone. Although a variety of module sizes are available, the one suitable for many semiconductor applications is 10 cm in diameter, 80 cm in length and contains about 100 ®bers housed in a PVDF shell. Each ®ber has an inside diameter of 1.7 mm, a wall thickness of 0.5 mm, and a pore size of 0.003 mm. The ®bers are arranged as a helix, a geometry which offers higher shell side mass

transfer coef®cients than one with ®bers parallel to the shell. Furthermore, the nature of expanded PTFE allows each ®ber to serve as a point-of-use particle ®lter, an important advantage in semiconductor manufacturing. Gas and liquid ¯ow rates of 3 and 10± 20 l minÿ1 are typical, with the gas on the tube side [142]. Pall Corporation (East Hills, NY) offers the SeparelTM (a trademark of Dainippon Ink and Chemicals) EFM-530 module for use in ultrapure water applications (see Section 6.10). The product is made from non-porous polyole®n ®bers that are woven into a fabric and wrapped around a central core. Water introduced through the central core ¯ows normal to the ®bers, which are either maintained under vacuum, ¯ushed with nitrogen sweep gas, or both. The shell is made from a clean type of PVC known as EslocleanTM, but PVDF will also be available in the near future. The non-porous membrane sets this product apart from the others discussed in this section, which all use microporous membranes. Non-porous membranes give higher selectivity of oxygen and other dissolved gases over water [66]. Hollow ®ber modules designed for pressure-driven ®ltration are generally con®gured as parallel ¯ow devices such as the contactor shown in Fig. 2, where the two ¯uids ¯ow parallel to each other on opposite sides of the membrane. In principle, there is no reason why these ®ltration modules cannot be used for concentration-driven ¯uid/¯uid mass transfer, where their simplicity (hence ease of manufacture) would be an

A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

65

Fig. 2. A parallel flow hollow fiber module (redrawn from [137]).

important advantage. On the other hand, con®gurations such as those discussed in the preceding paragraphs offer substantially higher mass transfer rates; this topic is explored in more detail below (Sections 5.2.3 and 5.2.4). Commercially available ®ltration modules which would be suitable for concentration-driven processes are described in Table 1. 3. Theoretical considerations

The schematic in Fig. 3 depicts mass transfer for a typical membrane contactor in liquid/liquid extraction service. Here the membrane is hydrophobic, so that the pores are wetted by the organic ¯uid. The solute of interest encounters three resistances in series upon transfer from the aqueous to the organic phase: the aqueous phase boundary layer, the membrane and the organic phase boundary layer [58]. The concentration pro®le is discontinuous at the interface according to the equilibrium relationship (1)

Here Cio and Ciw are the solute concentrations in the organic and aqueous phases, respectively, and mi is the partition coef®cient. For a hydrophobic membrane with the organic phase on the shell side, the driving force based on the organic phase concentration, and no chemical reaction, the resistance-in-series equation is 1=Ko di ˆ 1=kos do ‡ 1=kmo dlm ‡ mi =kwt di :

1=Kw di ˆ 1=kwt di ‡ 1=mi kmo dlm ‡ 1=mi kos do :

(2)

(3)

For a hydrophilic membrane with aqueous ¯ow through the tubes and no chemical reaction, the equations are 1=Ko do ˆ 1=kos do ‡ mi =kmw dlm ‡ mi =kwt di

3.1. Resistance-in-series model

Cio ˆ mi Ciw :

Here Ko is the overall mass transfer coef®cient, kos, kmo and kwt are the individual mass transfer coef®cients, and do, dlm, and di are the outside, log mean and inside diameters, respectively. For the driving force based on the aqueous phase concentration,

(4)

for the driving force based on the organic phase concentration, and 1=Kw do ˆ 1=kwt di ‡ 1=kmw dlm ‡ 1=mi kos do

(5)

for the driving force based on the aqueous phase concentration. The terms on the right-hand side of Eqs. (2), (3), (4) and (5) correspond to the resistances of the two boundary layers and the membrane. Analogous equations can be written for the organic phase on the tube side, and corresponding sets can be written for gas absorption/stripping, with the partition coef®cient mi replaced by the Henry's law constant. The complete sets of equations are given by Prasad and Sirkar [93] and Sirkar [122] for liquid/liquid and gas/ liquid contactors, respectively. Eqs. (2), (3), (4) and (5) assume the following [96]:  The system is at steady state.  Equilibrium exists at the fluid/fluid interface.  Pore size and wetting characteristics are uniform throughout the membrane.  The curvature of the fluid/fluid interface does not significantly affect the rate of mass transfer, the

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Manufacturer

A/G Technology (Needham, MA) Koch Membrane Systems (Wilmington, MA) Microdyn Technologies (Wuppertal, Germany) Millipore (New Bedford, MA) a

Materials of construction

Fiber diameter (mm)

Surface area (m2)

Pore size

Module length (cm)

Fiber

Housing

Potting

Polysulfone

Polysulfone

Epoxy

0.25±3

0.0015±28

1000 NMWCb-0.65 mm

18.5±120

Polysulfone, polyacrylonitrile, inorganic carbon Polypropylene, sulfonated polyether sulfone, polyethylene, regenerated cellulose Polysulfonee

PVC, polysulfone, 316LSSc, ARMYLOR1d Polypropylene, 316LSS

Epoxy

0.5±3.2

0.019±69.7

1000 NMWC-0.2 mm

17.8±182.9

Polyurethane, polypropylene

0.2±5.5

0.02±25

10 000 NMWC-0.4 mm

25±304.9

Polysulfone

Epoxy, polyurethane

0.5±1.1

0.03±5

3000 NMWC-0.1 mm

63.8±109.2

These are commercially available filtration modules which could be used for concentration-driven mass transfer; see text (Section 2). Information in this table was obtained from conversations with the manufacturers or from their literature. b Nominal molecular weight cutoff. c Type 316L stainless steel. d A PTFE-lined steel. e With polypropylene fiber wrap.

A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

Table 1 Characteristics of commercially available parallel flow hollow fiber modulesa

A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

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Fig. 3. Resistance-in-series model for a hydrophobic membrane.

equilibrium solute distribution or the interfacial area.  No bulk flow correction is necessary, i.e., mass transfer is described adequately by simple filmtype mass transfer coefficients.  No solute transport occurs through the non-porous parts of the membrane.  The two fluids are virtually insoluble in each other.  The equilibrium solute distribution is constant over the concentration range of interest. It is clear from Eqs. (2), (3), (4) and (5) that high and low values of the partition coef®cient mi result in a controlling resistance. For example, consider the extraction of a high mi solute with a hydrophobic membrane, where the organic phase ®lls the pores. We see from Eq. (3) that for high values of mi, both the organic boundary layer and the membrane resistances become negligible, so that the aqueous boundary layer resistance controls. For a hydrophilic membrane, where the aqueous phase ®lls the pores, we see from Eq. (5) that both the aqueous boundary layer and the membrane resistances are important. On the other hand, with a low mi solute and a hydrophobic mem-

brane, Eq. (2) indicates that both the organic boundary layer and the membrane resistances are signi®cant. From Eq. (4), only the organic phase boundary layer resistance is important if a hydrophilic membrane is used. These observations make sense intuitively: a high mi solute has a higher af®nity for the organic phase, so that one expects minimal mass transfer resistance there. Similarly, minimal resistance is expected through the aqueous boundary layer for low mi solutes since they have a higher af®nity for water. The results of D'Elia et al. [34] illustrates this point nicely. These workers used a hydrophobic membrane to study the extraction of (1) p-nitrophenol from water into amyl acetate, and (2) acetic acid from water into methyl amyl ketone with the organic phase on the shell side in each case. They found that in the ®rst case the aqueous boundary layer resistance controlled, while in the second case both the organic layer and the membrane resistances were important. The results follow from Eqs. (3) and (2) along with the mi values, which were 70 (high) and 0.4 (low) for the nitrophenol and acetic acid extractions, respectively.

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A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

In general, the membrane resistance is important. However, as discussed in more detail below, usually this additional resistance is more than offset by the higher interfacial area compared to conventional contactors. 3.2. Breakthrough pressure In membrane contactor operation the pores are ®lled with the ¯uid that wets the membrane. Nonpolar ¯uids will wet a hydrophobic membrane such as polypropylene, while polar ¯uids will wet a hydrophilic one such as polyacrylonitrile. To prevent the wetting ¯uid from permeating into the non-wetting ¯uid, the pressure of the latter is maintained slightly higher. Furthermore, the non-wetting ¯uid does not displace the ¯uid in the pores as long as the pressure on the non-wetting side is kept below some critical value known as the breakthrough pressure [61,62,96]. For liquid/liquid extraction the breakthrough pressure is given by [102]:
p ˆ …2 cos †=r;

(6a)

while for gas/liquid contacting the correct equation is
p ˆ …2 cos †=r:

(6b)

Here is the interfacial tension,  the surface tension,  the contact angle and r is the pore radius. The contact angle  is the angle formed between the wetting ¯uid and the membrane pore, as measured in the wetting ¯uid [82]. This angle increases with increasing polarity difference between the membrane material and the ¯uid. For example, water formed a contact angle of 1008 with an 80/20 copolymer of tetra¯uoroethylene and chlorotri¯uoroethylene at 208C, while the contact angle formed by benzene and this hydrophobic polymer was only 458 [39]. From Eqs. (6a) and (6b), a higher contact angle leads to a lower breakthrough pressure, which makes sense intuitively. That is, since the higher contact angle is indicative of a higher polarity difference between the ¯uid and the membrane, one would expect that less energy would be needed to remove the wetting ¯uid from the pore, hence the lower breakthrough pressure. The effect of interfacial tension (or surface tension ) also makes sense. As the interfacial tension increases, more energy is required to bring the two ¯uids together, hence more pressure is needed to

displace one ¯uid from the pores with the other. The role of the pore radius r is related to the ratio of the pore surface to its volume. This ratio increases as the pore radius decreases; the importance of the solid/¯uid interaction increases accordingly, resulting in a higher breakthrough pressure. Prasad and Sirkar [96] tabulated values of breakthrough pressure for a variety of membrane types and chemical systems; values range from about 100 to 500 kPa. Reed et al. [102] presented a table describing the relative magnitude of breakthrough pressure for various con®gurations of gas/liquid and liquid/liquid systems; factors considered include membrane type (hydrophobic or hydrophilic), ¯uid polarity (e.g., aqueous, polar organic, non-polar organic), and the wetting ¯uid (e.g., gas, aqueous, organic). A troublesome reduction in breakthrough pressure can occur in the presence of even trace levels of surfactant, which cause a reduction in interfacial or surface tension. Breakthrough can also be a problem with solvent systems where the interfacial tension is already low (e.g., water/n-butanol), or with a solvent that behaves like a surfactant (e.g., the pharmaceutical compound MK-819, mevinolinic acid) [96,102]. Three approaches to this problem have been suggested:  Coat the microporous fibers with a thin, highly permeable polymer or gel. Such a layer effectively addresses the breakthrough issue by stopping flow through the pores, yet it has no significant effect on diffusion [35,102].  Add a compound (e.g., isopropanol) which increases the interfacial tension [94].  Use a smaller pore size, which increases the breakthrough pressure in accordance with Eqs. (6a) and (6b). However, small pores can lead to hindered diffusion of large molecules such as proteins, causing a substantial reduction in mass transfer rate [96]. The majority of work on microporous hollow ®ber membrane contactors reported to date has been with membranes having a uniform pore diameter across the entire membrane thickness. Such membranes have a characteristic breakthrough pressure for a given chemical system, given by Eqs. (6a) and (6b). However, some work with asymmetric membranes has been reported, e.g., Sirkar et al. [127] studied solvent extraction using microporous hydrophilic alumina

A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

Ê on the inside tubes with pore diameters of 40±100 A Ê and up to 100 000 A on the outside. Such a membrane has no single breakthrough pressure; instead the breakthrough pressure varies with position as the pore radius changes, in accordance with Eqs. (6a) and (6b). By exceeding the breakthrough pressure on the side of the membrane with the smaller pore radius but not the larger, one can immobilize the interface within the pore rather than that at its mouth; such an approach has advantages in certain situations:  If the membrane resistance is significant, the interface can be shifted so that more of the pore volume is occupied by the fluid offering the lower resistance to mass transfer.  As mentioned above, hindered diffusion of large molecules can occur with small pores. In this case, the small pore radius on one side of an asymmetric membrane provides a workable breakthrough pressure, while the increasing pore radius across the membrane thickness leads to a substantial reduction in hindered diffusion [60,96]. Unlike dispersed phase contactors, with membrane contactors the pressure difference between ¯uids generally has no signi®cant effect on the mass transfer coef®cient [91,93,96]. 4. Mass transfer performance: comparison with conventional equipment As with conventional contacting equipment, with membrane contactors the task of the design engineer is to determine the contactor length necessary to achieve the desired extent of mass transfer. The well-known equation for traditional multistage contacting devices also applies here: L ˆ HTU  NTU:

(7)

In the literature, LTU is sometimes used instead of HTU since the membrane contactor is not necessarily oriented vertically. NTU is de®ned in terms of the ¯uid ¯ow rates, equilibrium solute distribution (i.e., partition coef®cient or Henry's Law constant), inlet concentrations and outlet concentrations [133]. An expression for the length L can be obtained by performing a differential solute balance, and then integrating over the length of the contactor. When the

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results are combined with the de®ning expression for NTU, the following equation for HTU is obtained from Eq. (7): HTU ˆ v=Ka;

(8)

where v is the ¯uid velocity, K the overall mass transfer coef®cient, and a is the interfacial area available for mass transfer [96]. As in conventional mass transfer equipment, the HTU value can be based on either the organic or the aqueous phase concentration driving force. As mentioned earlier, membrane contactors offer substantially lower HTU values than conventional mass transfer devices. Interestingly, the enhanced ef®ciency is not the result of a higher mass transfer coef®cient; indeed the mass transfer coef®cients obtained with membrane contactors are usually about the same or even slightly lower than those obtained with traditional dispersive contactors. Instead the improved performance results from the plethora of interfacial area, which remains constant with changes in operating conditions or ¯uid physical properties. Membrane contactors offer up to 30 times more area than what is achievable in ordinary gas absorbers and 500 times what is obtainable in liquid/liquid extraction columns [28]. Qi and Cussler [98] explained that the HTUs obtained in conventional towers all tend towards the same value (within a factor of 10 or so). This observation can be explained by considering Eq. (8). Generally the mass transfer coef®cient increases with increasing ¯ow rate (i.e., v), and the interfacial area decreases with decreasing ¯ow rate; as a result, the HTU remains about the same over a wide range of v. On the other hand, with membrane contactors the mass transfer coef®cient drops slightly at low ¯ow rates but the interfacial area remains unchanged, leading to very low values of HTU. Several authors have compared mass transfer performance of membrane contactors to that of conventional equipment. For example, Ding et al. [35] determined the mass transfer coef®cient obtained with a hollow ®ber contactor in the extraction of d-leucine from a racemic aqueous mixture into a 1-octanol solution of N-n-dodecyl-l-hydroxyproline, then compared their results to those reported by others for dispersed phase contactors. The ka values were 0.053, 0.0007 and 0.00005 sÿ1 for the membrane

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A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

contactor, a high-ef®ciency rotating column [132] and a conventional extractor [133], respectively. Sims et al. [121] used a membrane contactor to extract orange aroma into near-critical carbon dioxide, then compared their results to those obtained by Schultz et al. [107] with a Schiebel column. The extract from the membrane contactor contained substantially higher levels of key components such as ethyl butyrate, ethyl acetate, octanal, terpinen-4-ol and citronellal, and the membrane contactor was reportedly an order of magnitude more ef®cient. Similarly, Seibert and Fair [108] found that the extraction of hexanol from water into octanol was 10 times more ef®cient with a hollow ®ber membrane contactor than with a column containing type 2 structured packing; the authors attributed the higher ef®ciency to the signi®cantly greater mass transfer area. Reed et al. [102] gave the following values of interfacial area per unit volume (ftÿ1) for various types of contactors:  Free dispersion columns: 1±10.  Packed/trayed columns: 10±100.  Mechanically agitated columns: 50±150.  Membranes: 500±2000. Prasad and Sirkar [93] reported HTU values as a function of ¯ow rate for liquid/liquid extractions with chemical systems covering a range of partition coef®cients (0.013±50) for both hydrophilic and hydrophobic membranes. Values as low as 23 cm were obtained, and HTU was found to be independent of interfacial tension. Similarly, Prasad and Sirkar [95] reported HTU values of 3±15 cm in the simultaneous extraction of 4-methylthiazole (MT) and 4-cyanothiazole (CNT) from an actual process stream provided by Merck (Rahway, NJ), using a hydrophobic membrane and either toluene or benzene as the extraction solvent. As explained by the authors, these HTU values were probably lower than any values reported in the literature on packed and agitated columns in solvent extraction service. The closest literature value reported was 17 cm, obtained by Steiner and Hartland [130] in their work with an enhanced coalescence plate agitated column operated at high agitator speed (250 rpm). Prasad and Sirkar [94] compared the cost of hollow ®ber membrane contactors and conventional extractors for the extraction of the pharmaceutical compound mevinolinic acid or MK-819. The compound is

extracted from isopropyl acetate into water at high pH, then back-extracted into isopropyl acetate at low pH. For the high pH extraction, the cost of a centrifugal extractor (e.g., a Podbielniak) would be 3±20 times higher than that of a membrane contactor for a membrane life of 0.5±3 years. Similarly, for the low pH extraction the cost of a mixer±settler would also be 3± 20 times higher. This work is discussed in more detail in Section 6.7. Kreulen et al. [64] compared the performance of a bubble column to that of a microporous polypropylene hollow ®ber membrane contactor for the absorption of oxygen into glycerol/water solutions of various concentrations. The membrane module (interfacial area: 20 cm2/cm3; ®ber length: 18 cm) had the same liquid holdup and liquid height as the bubble column (diameter: 4 cm; liquid holdup: 200 cm3). In light of the other examples cited in this section, it is somewhat surprising that the kLa for the bubble column was actually higher than that of the membrane contactor at low glycerol concentrations. For example, with water alone the kLa was 0.20 sÿ1 for the bubble column vs. 0.023 sÿ1 for the membrane contactor. Both kLa values decreased with increasing glycerol concentration, but the value for the bubble column decreased faster. Values were higher for the membrane contactor at high glycerol concentrations, e.g., 0.002 and 0.0012 sÿ1 for the membrane contactor and the bubble column, respectively, at 80% glycerol. The authors explained that the increasing viscosity with increasing glycerol concentration led to a reduction in the mass transfer coef®cient for both devices; however, the interfacial area decreased in the bubble column while it stayed the same in the membrane contactor, hence the slower degradation in kLa with the latter. Interestingly, when compared on the basis of degree of saturation (i.e., approach to equilibrium), the membrane contactor gave superior performance over the entire range of glycerol concentrations studied. For the bubble column the degree of saturation ranged from 84% at 0% glycerol to 20% at 80% glycerol, while the membrane contactor provided saturations of 87±70% over the same concentration range. The authors explained these results in terms of the type of gas/ liquid mixing: the bubble column behaved as a stirred tank while the membrane contactor was better described by plug ¯ow.

A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

Jansen et al. [52] looked at using a hydrophobic membrane contactor for CO2 removal from ¯ue gas in applications like ammonia production and natural gas separation. They proposed a design for 70% CO2 recovery from 600 m3/s of gas using a reactive solvent such as monoethanolamine to minimize the liquid side resistance to mass transfer. The authors assumed that diffusion through the gas-®lled pores controlled (see Section 5.3) and estimated the mass transfer coef®cient to be 0.02 m/s for a hollow ®ber with an outside diameter of 1 mm. The required membrane area turned out to be 35 500 m2; at a membrane cost of 50 D¯ (about $26)/m2, the equipment cost would be 30% lower than that of conventional gas absorption, and the total required investment would be 10% lower. Furthermore, the estimated pressure drop for the membrane contactor was half that of the conventional absorber, which would lead to a signi®cant reduction in energy cost. The authors pointed out that detailed engineering design would be required to con®rm these estimates. Additionally, problems with amine-based solvents wetting the membrane pores have been encountered by others studying this application [126]. Jansen et al. [52] also made a more general comparison of gas absorption with a membrane contactor vs. a packed column, both using a reactive solvent. Mass transfer coef®cients for the packed column were estimated from the correlation proposed by Onda et al. [81]. For the membrane contactor, the membrane mass transfer coef®cient was taken to be 0.03 m/s based on the authors' experimental work with commercially available membranes, and the gas side coef®cient was estimated from Eq. (31); the liquid phase resistance was assumed to be negligible for this reactive system (see Section 5.3). The inside and outside ®ber diameters were assumed to be 0.6 and 1.0 mm, respectively, and the speci®c surface area of the packing was taken as 108 m2/m3 (typical for a packing diameter of 2 in.). At a gas velocity of 3 m/s, the overall mass transfer coef®cient for the packed column was 1/3 or 1/5 of the value obtained for the membrane contactor, assuming a membrane speci®c surface area of 470 or 750 m2/m3, respectively. The improvement increased with decreasing gas velocity. Karoor and Sirkar [56] used a microporous polypropylene membrane contactor to study the absorption of pure CO2, pure SO2, CO2 from CO2/N2 mixtures,

71

and SO2 from SO2/air mixtures into water. They obtained mass transfer coef®cients ®ve and 10 times higher than those typically obtained in packed towers for absorption of CO2 and SO2, respectively. This work is discussed in more detail in Section 6.2.

5. Design considerations 5.1. Experimental determination of mass transfer coefficients Overall and individual mass transfer coef®cients can be determined experimentally according to the resistance-in-series model discussed above. While the approach described in the text that follows is for a liquid/liquid contactor, the procedure for a gas/liquid contactor is completely analogous if the partition coef®cient mi is replaced by the Henry's Law constant. First the overall mass transfer coef®cient is determined using the following form of the familiar ¯ux equation [93]: Qo Cio1 ˆ Ko a
Cilm :

(9)

Here Cio1 is the concentration of component i in the organic phase leaving the module; the inlet concentration is assumed to be zero. Note that Ko is the average rather than the local mass transfer coef®cient; also, the organic phase ¯ow rate Qo can be taken as constant because the solute concentration is assumed to be low. The log mean driving force in Eq. (9) is given by
Cilm ˆ …
Ci1 ÿ
Ci0 †=ln…
Ci1 =
Ci0 †:

(10)

The subscripts 1 and 0 refer to the respective ends of the module, and
Ci ˆ mi Ciw ÿ Cio :

(11)

The distribution coef®cient mi is either obtained from the literature or determined experimentally. Typically the feed concentration and the interfacial area are known, so that measurement of the ¯ow rate and outlet concentrations allows the mass transfer coef®cient to be calculated. Individual mass transfer coef®cients are determined next. For unhindered diffusion through organic ¯uid®lled pores, Prasad and Sirkar [93] gave the following

72

A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

equation for the membrane mass transfer coef®cient: kmo ˆ Dio M =‰M …do ÿ di †=2Š:

(12)

Here the porosity M, the tortuosity  M and the diameters are typically available from the membrane manufacturer. Porosities range from 0.2 to 0.9 but generally fall around 0.35, while tortuosities are typically 2 to 3. Although the pore size does not appear explicitly in this equation, it appears implicitly through the diffusivity. However, the effect of diffusivity is generally minor, so the pore size is usually unimportant as long as the diffusion is unhindered; the exception is gas-®lled pores where Knudsen diffusion is important [102]. Prasad and Sirkar [93] gave a similar equation for aqueous ¯uid-®lled pores. As an alternative to Eq. (12), the Wilson plot method [145] can be used to determine the membrane resistance, and also to gain insight into the effect of ¯uid velocities on individual mass transfer coef®cients [93]. Here we assume that the individual coef®cients for the organic and aqueous boundary layers are

proportional to v, where  is an empirical constant. In this case a plot of kÿ1 vs. vÿ gives a straight line; such a plot is known as a Wilson plot. For a hydrophobic membrane a low mi solute and/or a high aqueous ¯ow rate is used, so that the last term on the right-hand side of Eq. (2) becomes negligible. The overall mass transfer coef®cient is determined as a function of vo using Eqs. (9)±(11), then the Wilson plot of 1/Ko vs. 1=vo o is prepared. Here the value of o is selected as the one that provides the best straight line through the data points; the membrane resistance is given by the intercept of the Wilson plot. The approach is analogous for a hydrophilic membrane. Fig. 4 shows the Wilson plot presented by Prasad and Sirkar [93] for extraction of acetic acid from water into xylene using a hydrophobic membrane with the organic solvent on the shell side. The boundary layer resistances can be determined once the overall and membrane resistances are known. Consider a hydrophobic membrane and a low mi solute; to determine the organic boundary layer resis-

Fig. 4. Wilson plot for extraction of acetic acid from water into xylene using a hydrophobic hollow fiber module with the solvent on the shell side (data of [93]).

A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

73

tance, again the ¯ow rate of the aqueous phase is maintained at a high value. Since this eliminates the aqueous phase resistance of Eq. (2), the overall resistance is now just the sum of the organic boundary layer and the membrane resistances. The overall resistance is determined in the manner described above (Eqs. (9)±(11)), then the membrane resistance (known from Eq. (12) or the Wilson plot) is subtracted to obtain the organic boundary layer resistance. The aqueous phase resistance is determined in the same manner except the organic rather the aqueous phase ¯ow rate is maintained at a high value. The approach is similar for a high mi solute (Eq. (3)) or a hydrophilic membrane (Eqs. (4) and (5)). Cooney and Poufos [22] suggested a different approach. These researchers determined individual resistances in the extraction of (a) phenol from water into n-octanol, and (b) phenol from hexane into water, using a hydrophilic membrane with the organic phase on the shell side in both cases. Unlike most other work cited in this review, the extracting solvent did not ¯ow, i.e., it was held stagnant. The proposed equation for the mass transfer coef®cient,

mass transfer coef®cient was again measured as a function of tube side ¯ow rate, then the corresponding membrane plus aqueous boundary layer resistances (determined from the short run) were subtracted to obtain the shell side resistances. In this manner Cooney and Poufos [22] showed that the organic phase boundary layer resistance was unimportant in the phenol/water vs. n-octanol system, yet it was in fact signi®cant for the phenol/hexane vs. water system. These results are consistent with Eqs. (4) and (5) and the mi values for phenol/water vs. n-octanol and phenol/hexane vs. water which are 396 and 0.96, respectively.

Ka ˆ …Qf =SL†ln…Cin =Cout †

Sh / Re Sc f …geometry†:

(13)

was derived by writing a mass balance across a differential section of the membrane contactor, then integrating over the length of the contactor. Two types of runs were performed: short and long. The short runs did not allow enough time for a signi®cant amount of solute to enter the extracting solvent, so that the solvent boundary layer resistance was eliminated. The sum of the other resistances might be expected to vary linearly with 1/Q1/3 in analogy with the equation of Sieder and Tate [120] for heat transfer: Nu ˆ 1:86…Re Pr d=L†1=3 …b =w †0:14 :

(14)

However, Cussler [27] and others have pointed out that Eq. (14) is not valid for very low Reynolds numbers, presumably due to free convection effects. In this range the appropriate plot is 1/Ka vs. 1/Q, and in fact Cooney and Poufos [22] obtained the expected straight line upon plotting their data in this manner. As explained above, the intercept of such a plot gives the membrane resistance. The organic phase resistance was determined by performing long runs to allow the solute concentration on the organic side to become non-zero. The overall

5.2. Prediction of individual mass transfer coefficients Correlations describing mass transfer coef®cients are important in the design of mass transfer equipment, and membrane contactors are no exception. For mass transfer devices in general and membrane contactors in particular, mass transfer coef®cients can be predicted using correlations of the form (15)

Here Sh, Re and Sc are the Sherwood number, the Reynolds number and the Schmidt number, respectively, and f is some function of geometry. Several authors have presented summaries of applicable correlations for both tube and shell side ¯ow [24,29,53,96,102,122,140]. A summary of equations presented in this review is given in Table 2. 5.2.1. Tube side flow The following correlation has been shown by several investigators [140,146] to predict tube side mass transfer coef®cients with reasonable accuracy: Sh ˆ 1:62…d 2 v=LD†1=3 :

(16)

This is the well-known LeÂveÃque solution, a limiting case of the more general Graetz solution applicable when the Graetz number is large [46,47,63,71, 102,139]. Note that Eq. (14) is a variation. Eq. (16) overestimates experimentally determined mass transfer coef®cients at low ¯ows; this can be attributed to non-uniform ¯ow caused by polydispersity in hollow ®ber diameter [140].

74

Table 2 Summary of equations for predicting individual mass transfer coefficients Equation no.

Equation

Comments

Reference

Tube side flow 16 Shˆ1.62(d2v/LD)1/3

[71]

17

[140]

kmoˆDioM/[ M(doÿdi)/2]

Shell side flow parallel to the fibers 19 Shˆ [de(1ÿ)/L]Re0.6 Sc0.33 20 Shˆ1.25(Re de/L)0.93Sc0.33 21 Shˆ0.019Gz 22 Shˆ8(Re de/L)Sc0.33 26

Shˆ(0.53ÿ0.58)Re0.53 Sc0.33

Shell side flow across the fibers 28a Shˆ0.15Re0.8 Sc0.33 28b Shˆ0.12Re Sc0.33 30a Shˆ6.0Gz0.35 30b Shˆ1.25Gz 31 Shˆ0.9Re0.5 Sc0.33 32a 32b 33 34 35

Shˆ1.38Re0.34 Sc0.33 Shˆ0.90Re0.40 Sc0.33 Shˆ0.61Re0.363 Sc0.333 Shˆ1.45Re0.32 Sc0.33 Shˆ0.24(Re de/L)0.59Sc0.33

Hollow fiber fabric 36 Shˆ0.82Re0.49 Sc0.33 38 39

Shˆ0.18Re0.86 Sc0.33 Shˆ0.46Re0.46 Sc0.33

40

Shˆ0.57Re0.31 Sc0.33

Assumes unhindered diffusion through organic fluid-filled pores. An analogous expression can be written for pores filled with aqueous fluid

[93]

is 5.8 for hydrophobic and 6.1 for hydrophilic fibers; 0
[93] [146] [140] [29]

Developed using a cylindrical tube, a helically wound bundle and a rectangular bed of fibers, with Re>2.5 Same geometries as Eq. (28a) except Re<2.5; results attributed to polydisperse channels between fibers Crimped flat membrane, Gz>11 Crimped flat membrane, Gz<11 Proposed for gas absorption with a reactive solvent using a rectangular module; 1
[140] [140] [140] [140] [52]

[24]

[146] [146] [26] [3] [53]

Average of correlations developed using four different configurations, with 0.01
[141]

[13]

Fast reaction with a low solute concentration in the liquid Instantaneous reaction with negligible gas phase resistance

[70] [70]

[137] [137]

a

Gas/liquid mass transfer with chemical reaction 42 kˆ(kRDAL)1/2 43 k0 ˆk(1‡DBLCB/bDALCAi) a

Analogous equations can be written for liquid/liquid mass transfer.

A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

Membrane 12

In general, predicts tube side mass transfer coefficients with reasonable accuracy for Gz>4 but overestimates them for Gz<4 hShi ˆ Sh‰1 ÿ …18Sh=Gz ‡ 7Š20 ‡


Š Accounts for polydispersity of hollow fiber diameters. hShi is the average Sherwood number; Sh is the Sherwood number expected for a uniform distribution of fiber radii

A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

Indeed, Park and Chang [84] showed that tube side ¯ow distribution is often not uniform. Using highspeed photography and dye tracer studies, these researchers determined that the distribution depends on the inlet manifold type (cylindrical or conical), manifold height, tube length, ®ber inner diameter, shell diameter, ®ber packing density and Reynolds number. Nearly uniform ¯ow was achieved for certain shell and tube geometries with long manifolds at low Reynolds numbers; however, non-uniform ¯ow was the rule rather than the exception. Experimental data agreed fairly well with model predictions. Wickramasinghe et al. [140] assumed a Gaussian distribution of ®ber radii, then integrated the basic equation for the mass transfer coef®cient over all radii and obtained the following power series solution for the average tube side coef®cient: hki ˆ k‰1 ÿ …9kV=QR0 ‡ 7†20 ‡


Š:

(17)

Here k is the mass transfer coef®cient expected for a uniform distribution of ®ber radii, V the average volume occupied by one ®ber, R0 the average ®ber radius, and 0 is the standard deviation of ®ber radii divided by the mean. Eq. (17) was developed for low ¯ow conditions, i.e., the Graetz number less than four; at these low ¯ows, the quantity in brackets is less than one. Since the distribution of ®ber diameters leads to low ¯ows in some ®bers, we see that polydispersity can lead to a reduction in the mass transfer coef®cient. The hypothesis was veri®ed experimentally upon stripping oxygen from water into water-saturated nitrogen using microporous polypropylene hollow ®bers. Correlations inferred from the experimental data agreed with Eq. (16) at high ¯ow rates, but gave lower mass transfer coef®cients at low ¯ows in accordance with Eq. (17). 5.2.2. Shell side flow parallel to the fibers In contrast to tube side ¯ow, shell side ¯ow is not clearly understood. A number of correlations like Eq. (15) have been proposed, but none is applicable to a wide range of systems in the manner Eq. (16) is generally applicable to tube side ¯ow. Knudsen and Katz [63] proposed the following equation for shell side heat transfer in a shell-andtube exchanger: Sh ˆ 0:022Re0:6 Sc0:33 :

(18)

75

In general this correlation does not predict available experimental data on mass transfer in membrane contactors, probably because of the lack of dependence on geometry [139]. However, variations of this equation that include geometrical dependence have been used successfully; examples are cited in the discussion that follows. Prasad and Sirkar [93] found that an equation similar to Eq. (18) correlated their data on liquid/ liquid extraction with various solutes and solvents: Sh ˆ ‰de …1 ÿ †=LŠRe0:6 Sc0:33 :

(19)

Here  is the packing fraction. The constant was found to be 5.8 and 6.1 for hydrophobic and hydrophilic membranes, respectively. Yang and Cussler [146] obtained the following correlation in their work with gas absorption and stripping: Sh ˆ 1:25…Re de =L†0:93 Sc0:33 :

(20)

The authors had no explanation for the unexpectedly high Reynolds number dependence vs. available correlations for heat transfer (e.g., Eq. (18)). However, in practice Eqs. (19) and (20) give similar results, even though the latter has a higher Reynolds number dependence and does not contain the packing fraction  [139]. Moreover, other studies have also revealed a linear (or nearly so) dependence on ¯ow. For example, Wickramasinghe et al. [140] obtained the following correlation in their work with gas absorption and stripping using close-packed ®bers: Sh ˆ 0:019Gz:

(21)

These researchers argued that correlating the Sherwood number with the Graetz number instead of the product of the Reynolds and Schmidt numbers makes more sense. Their argument assumed that the Sherwood number changes little with changes in viscosity, which they explained is true if the velocity pro®le is established quickly. Dahuron and Cussler [29] correlated their data on extraction of proteins as follows: Sh ˆ 8…Re de =L†Sc0:33 :

(22)

These investigators explained that the deviation from the expected Reynolds number exponent of 0.6 may have been due to channeling around the

76

A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

®bers; they also pointed out that their results depended on the accuracy of the diffusion coef®cients of the proteins they used, and that they had no direct measure of these coef®cients. The work of Dahuron and Cussler is discussed in more detail in Section 6.6. Lemanski and Lipscomb [69] looked at the effect of shell side residence time distribution (RTD) on mass transfer performance. Plug ¯ow would be obtained with an ideal contactor, but in real shell side ¯ow the need to distribute ¯uid across the ®ber bundle as well as along it tends to broaden the RTD. Mass transfer is more ef®cient for ¯uid elements with a shorter residence time because the concentration driving force remains higher than that for elements which spend more time in the shell. Furthermore, since the mass transfer coef®cient is a function of the local ¯uid velocity, higher local mass transfer coef®cients are obtained for ¯uid elements in cross¯ow than for those ¯owing parallel to the ®bers. Lemanski and Lipscomb [69] characterized shell side mass transfer in terms of the ratio of the mass transfer coef®cient for plug ¯ow (at the volume average velocity) to the actual, intrinsic mass transfer coef®cient. They showed that the ratio depends on two dimensionless groups: the module geometry number MG ˆ …rr =zz †…L=R†2

(23)

and the mass transfer number MT ˆ ka:

(24)

Here  is the Darcy permeability tensor and  is the space time. By volume averaging the pertinent conservation of mass and momentum equations, the ratio was calculated over a range of values of MT and MG; results are shown in Fig. 5. At small values of MT, performance was superior for lower values of MG; however, as MT increased, the performance decreased substantially for lower MG values but changed little for higher ones. The authors concluded that commercial modules should have an MG value of at least 10 to minimize the effect of MT on performance. Seibert et al. [109] studied shell side bypassing in the extraction of acetone from toluene into water using a hydrophobic hollow ®ber module. They found from dye tracer studies that the actual residence time was less than ideal by a factor of 3±4, and the extent of bypassing changed with ¯uid velocity according to the

Fig. 5. Dependence of kE/k on MT for various values of MG. Diamond: 1; filled circle: 2; triangle: 5; square: 75; open circle: 150 (redrawn from [69]).

following empirical equation:  ˆ exp…ÿCvs †:

(25)

Here  is the fraction of ¯uid that bypasses the hollow ®bers, C is an empirical constant, and vs is the shell side ¯uid velocity. The value of C increased with decreasing ®ber packing density; the authors reported values of 0.38, 1.4 and 2.8 s/cm for ®ber counts of 7500, 3000 and 1500, respectively. Other investigators have also studied the effect of ®ber packing density on shell side mass transfer. For example, Wickramasinghe et al. [140] looked at the absorption of oxygen into water with the latter on the shell side using a module containing 12 000 or 7500 ®bers. Mass transfer coef®cients for the more loosely packed module were three times higher; the authors attributed the difference to channeling. Yang and Cussler [146] also reported that the mass transfer coef®cient was lower for closely vs. loosely packed ®bers in the absorption of oxygen into water, and there was no change in the mass transfer coef®cient with Reynolds number with the former. They attributed these results to major channeling through the closely packed ®bers so that mass transfer was controlled by diffusion through a stagnant ®lm between the ®bers, hence the lack of Reynolds number dependence. Costello et al. [24] looked at the effect of ®ber packing density on mass transfer upon stripping of oxygen from water into nitrogen gas, using a hydrophobic membrane with water and nitrogen ¯owing on

A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

the shell and tube sides, respectively. They pointed out that most of the published correlations for the shell side mass transfer coef®cient (e.g., Eqs. (19), (20) and (22)) were developed using contactors with low ®ber packing densities (3±40%). Values of  (Eq. (15)) for these correlations range from 0.6 to 1.0, characteristic of turbulent ¯ow. However, these correlations were developed for shell side Reynolds numbers in the laminar range; this suggests the presence of local turbulence as explained below. The correlations did not predict the experimental data obtained by Costello et al. at higher packing densities, i.e., 32±76%. Costello et al. [24] proposed the following equation for the shell side mass transfer coef®cient: Sh ˆ …0:53 ÿ 0:58†Re0:53 Sc0:33 :

(26)

Their data suggested values of the Re exponent ranging from 0.50 to 0.57, characteristic of entry region conditions, over a range of packing densities. According to the authors, these results were indicative of splitting and remixing of ¯uid streams, which constantly presented the membrane surface with fresh ¯uid. Unlike similar correlations proposed by others (e.g., Eqs. (19) and (20)), there is no dependence on d/L in Eq. (26), which implies that the entrance length is unrelated to the module length; instead it is probably related to stream splitting and recombining that occurs repeatedly along the length of the module. One can solve the differential mass and momentum balances to obtain a theoretical prediction of the effect of packing density on the mass transfer coef®cient. According to that prediction, the coef®cient increases steadily with increasing packing density until a maximum is reached at a density of 65±70%. The subsequent decline is caused by ``dead zones,'' which are regions where the ®bers are very close together or actually touch. For packing densities less than 65%, experimental coef®cients are substantially higher than expected based on the theory. The reason is that the theory assumes the ®bers are packed uniformly, which is not the case in real systems. This non-uniformity promotes turbulence, i.e., the stream splitting and recombining mentioned above, which in turn leads to mass transfer coef®cients that are higher than those theoretically possible with axial laminar ¯ow [24]. Rogers and Long [104] also attributed at least part of the dif®culty in predicting shell side mass transfer coef®cients to incorrect assumptions regarding shell

77

side ¯ow distribution. They explained that most models divide shell side ¯ow among ®nite grids whose size and shape are determined primarily by the geometry of the ®ber layout, which is assumed to be (but rarely is) uniform. Rogers and Long used a technique known as random sequential addition (RSA) to generate an ensemble of randomly oriented ®bers. The maximum packing density obtained, known as the jamming limit, was approximately equal to the packing density used in at least one commercially available module, i.e., the Liqui-Cel module from CELGARD LLC (Fig. 1). Interestingly, the placement of 2186 non-overlapping ®bers using RSA required 20 953 544 iterations and 80.81 h on a Pentium 100 MHz PC. With the random placement of ®bers in this manner, shell side ¯ow was divided into Voronoi cells or polygons, de®ned as the boundary around each tube where the momentum gradient is zero or the velocity is maximized. These boundaries were identi®ed using the software known as Voronoi tessellation code. Rogers and Long [104] modeled the performance of hollow ®ber membrane contactors when used for extraction accompanied by chemical reaction. Mass transfer coef®cients were calculated using published correlations, and shell side ¯ow maldistribution was estimated using RSA and Voronoi polygons as explained above. Model predictions were compared to experimental data for the extraction of uranyl nitrate from a nitric acid solution into tributyl phosphonate (TBP) in a hydrocarbon diluent. The extraction was accompanied by the reaction ÿ UO‡‡ 2 ‡ 2NO3 ‡ 2TBP $ UO2 …NO3 †2
2TBP

(27)

The enhancement of mass transfer by the reaction was included in the model by the use of a facilitation factor similar to the Thiele modulus. The authors concluded that the agreement between experimental data and model predictions was signi®cantly better than the predictions of lumped parameter models that assume uniform shell side ¯ow. Mass transfer with chemical reaction is discussed in more detail in Section 5.3 below. 5.2.3. Alternative hollow fiber geometries In the early days of hollow ®ber membrane contactors the controlling resistance was typically the membrane itself, primarily because membranes were

78

A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

thick and permeabilities were low. Since that time, our knowledge of membrane materials and manufacturing methods has grown dramatically, and membranes with much lower mass transfer resistance have been developed. This progress focused attention on the other mass transfer resistances encountered in parallel ¯ow contactors (see Fig. 2), particularly the shell side boundary layer, hence the work reviewed in the previous section. This altered focus has fostered the development of several alternative geometries that offer higher mass transfer coef®cients than conventional parallel ¯ow. For example, Wickramasinghe et al. [140] stripped oxygen from water into nitrogen using four module con®gurations, each with water ¯owing outside and across the microporous polypropylene hollow ®bers. The four geometries were a cylindrical tube bundle, a helically wound bundle, a rectangular bed of ®bers and a crimped ¯at membrane. Mass transfer coef®cients for the ®rst three con®gurations were described by Sh ˆ 0:15Re0:8 Sc0:33

(28a)

for Re>2.5, and Sh ˆ 0:12Re Sc0:33

(28b)

for Re<2.5. Sherwood numbers obtained at low ¯ow rates (Eq. (28b)) were lower than those predicted by the correlation applicable to high ¯ow rates (Eq. (28a)). The authors attributed this to uneven ¯ows caused by polydisperse channels between ®bers; they presented the following equation to account for the effect of polydispersity: hki ˆ k‰1 ÿ 9kV=QR0 …0 …1 ÿ 0 ††1=2 ‡ 7Š20 ‡


Š; (29) where 0 is the void fraction of ®bers in the module. This equation is analogous to Eq. (17) and was derived in a similar manner. Both assume a Gaussian distribution of ¯ow channel radii; for the tube side this assumption was supported by measurements of ®ber inside diameter, but for the shell side it remained speculative. Furthermore, on the tube side a larger ®ber diameter results in a larger ¯ow through that ®ber for the entire module length, but on the shell side the situation is not so clear. For these reasons the authors had more con®dence in the general applicability of Eq. (17) to tube side ¯ow than in the utility of Eq. (29)

for shell side ¯ow, even though the existence of polydisperse gaps seemed like a reasonable explanation for inferior module performance at low ¯ows. Wickramasinghe et al. [140] obtained the following correlations for the crimped ¯at membranes: Sh ˆ 6:0Gz0:35

(30a)

for Gz>11, and Sh ˆ 1:25Gz

(30b)

for Gz<11. The cube root dependence in Eq. (30a) is consistent with expected results for ¯ow in a slit. The linear dependence in Eq. (30b) is consistent with Eq. (28b). Wickramasinghe et al. [140] compared performance of the various module geometries studied in terms of oxygen removal ef®ciency. When performance was compared on the basis of equal ¯ow per unit membrane area, best results were obtained with the rectangular bundle, which stripped 98% of the incoming oxygen. The cylindrical bundle, helical bundle, and crimped ¯at plate removed 82%, 86% and 72%, respectively. All of these cross¯ow designs were substantially more ef®cient than a parallel ¯ow cylindrical module, which stripped only 7% of the oxygen in the incoming water. Findings were similar when module performance was compared on the basis of equal ¯ow per unit volume. For gas absorption applications, Jansen et al. [52] proposed a rectangular design with gas ¯ow outside and normal to the ®bers. According to the authors, such a design offers a number of advantages, including high mass transfer, low pressure drop, well-de®ned ¯ow conditions on both sides of the membrane, gas ¯ow unobstructed by ®ber potting, and ability to stack modules. A preliminary optimization study suggested that mass transfer was more ef®cient with an in-line rather than a staggered ®ber arrangement. Best results were obtained when the transversal pitch was twice the outside diameter, and the longitudinal pitch divided by the outside diameter was between 1.1 and 3.0; with this optimal geometry, Sh ˆ 0:9Re0:5 Sc0:33

(31)

for Re and Sc each between 1 and 1000, with Re based on the external ®ber diameter. Yang and Cussler [146] used hydrophobic hollow ®bers to study the removal of oxygen or CO2 from

A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

water into nitrogen. The water ¯owed outside and across the ®bers, which were housed in a conventional cylindrical shell. The authors obtained the following correlations for closely and loosely packed ®bers, respectively: Sh ˆ 1:38Re0:34 Sc0:33 ;

(32a)

Sh ˆ 0:90Re0:40 Sc0:33 :

(32b)

Both of these equations resemble published heat transfer correlations for single tubes rather than tube banks. While this is expected for loosely packed ®bers (Eq. (32b)), one would expect correlations for tube banks to predict data from closely packed ®bers more accurately, since the two are more similar geometrically. The authors were unable to explain the result. CoÃte et al. [26] obtained a similar correlation for loosely packed ®bers in experiments on oxygenation of water: Sh ˆ 0:61Re0:363 Sc0:333 :

(33)

Ahmed and Semmens [3] also developed a similar correlation in their work on bubbleless aeration of water using microporous polypropylene hollow ®bers: Sh ˆ 1:45Re0:32 Sc0:33 :

(34)

In related work, Johnson et al. [53] studied bubbleless aeration of water using a sealed hollow ®ber bundle uncon®ned in a jet stream. Their mass transfer correlation, Sh ˆ 0:24…Re de =L†0:59 Sc0:33

(35)

was veri®ed using different membrane lengths, jet ori®ce areas and jet liquid exit velocities. Bubble-free aeration is discussed in more detail in Section 6.8. Baf¯ed contactors (e.g., Fig. 1) were developed to overcome shortcomings of parallel ¯ow such as shell side bypassing. However, Seibert and Fair [108] found that bypassing does in fact occur with these modules in some cases. These researchers proposed a stage ef®ciency model that was an excellent predictor of mass transfer coef®cients for a shell side-controlled system (octanol/hexanol/water) and a mixed resistance system (butanol/succinic acid/water) at high shell side ¯ow rates but a poor predictor at low ¯ow rates. They suggested poor shell side distribution at low ¯ow rates as a possible explanation for the discrepancy between expected and observed results. According to the

79

authors, such problems are more severe in larger modules; little improvement is obtained by decreasing the ®ber density because ®bers tend to part and provide channels for bypassing. 5.2.4. Hollow fiber fabric Several of the investigators cited in the above discussion found that uneven ¯ow distribution led to a reduction in the mass transfer coef®cient. Along these lines, Cussler [28] pointed out that the inevitable non-uniform ®ber spacing in commercial modules results in non-uniform ¯ow and in turn lower mass transfer coef®cients than those obtained with handbuilt laboratory units, where precise spacing is more easily achieved. Cussler suggested the use of ®bers woven into a fabric to obtain more uniform spacing, and presented data showing that mass transfer coef®cients obtained using this approach were nearly as high as those with hand-built contactors. Wang and Cussler [137] and Wickramasinghe et al. [141] used modules employing a woven fabric to strip oxygen or toluene from water into nitrogen. Similarly, Bhaumik et al. [13] used hollow ®ber fabric to absorb CO2 from pure CO2 and from a CO2/N2 mixture into water. As explained below, all three groups obtained mass transfer coef®cients that were up to an order of magnitude higher than those achieved with non-woven ®bers. Wickramasinghe et al. [141] stripped oxygen from water using microporous polypropylene hollow ®bers arranged in one of the four con®gurations; in all cases the gas ¯owed on the tube side. The membrane modules used in this work were constructed by the authors, and not manufactured commercially. The ®rst con®guration (Fig. 6) was an annular bed of hollow ®bers wound helically around a central core; liquid entering the central core was forced radially outward by a plug so that the ¯ow was perpendicular to the ®bers. The second con®guration was similar to the ®rst except that the shell contained plugs and O-rings to provide multiple shell side passes. The third con®guration (Fig. 7) was similar to the second except that it was made with knitted hollow ®ber fabric rather than individual ®bers. Finally, in the fourth con®guration (Fig. 8), hollow ®ber fabric was mounted diagonally in an open-ended rectangular box, and the liquid entered through a tubular manifold. The whole box was towed through the liquid, forcing the liquid to ¯ow both through and around the box.

80

A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

Fig. 6. A module containing hollow fibers wound helically around a central core. Liquid entering the core is forced radially outward by the plug so that the flow is perpendicular to the fibers (redrawn from [141]).

Fig. 7. A module containing woven hollow fiber fabric wound helically around a central core. The plugs and O-rings provide multiple shell passes (redrawn from [141]).

Higher mass transfer coef®cients were obtained with the axially wound individual ®bers (i.e., the ®rst two con®gurations described above) than with modules manufactured commercially, especially at low ¯ows [140]. However, coef®cients for the axially wound modules were not as high as those obtained

with modules built one ®ber at a time [146]. These results support the argument that better mass transfer is achieved with more uniform ®ber spacing, since uniformity is usually best with hand-built and worst with commercially manufactured contactors; unfortunately, the bene®ts of hand-built modules are probably

Fig. 8. A module containing woven hollow fiber fabric mounted diagonally in an open-ended box (redrawn from [141]).

A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

not justi®ed by the higher manufacturing cost. Surprisingly, the mass transfer coef®cients for the unbaf¯ed (Fig. 6) and baf¯ed (the second con®guration described above) axially wound modules were about the same. Mass transfer coef®cients for the axially wound fabric (Fig. 7) were slightly higher than those obtained with the axially wound individual ®bers, yet still lower than those achieved with hand-built modules. When compared to performance of commercial modules [140], coef®cients were an order of magnitude higher with the fabric at low ¯ows, while at high ¯ows the improvement was not as dramatic. The authors attributed the superior performance of modules containing the fabric to higher uniformity of ®ber spacing compared to individual ®bers. The fabric is also easier to assemble than commercial modules, so that better mass transfer is achieved at a lower manufacturing cost. The mass transfer coef®cients for the rectangular module (Fig. 8) were similar to those obtained with hand-built modules [146], i.e., they were the highest of the four con®gurations studied. However, the complexity of these modules may outweigh their performance advantage. Wickramasinghe et al. [141] presented mass transfer correlations for hand-built modules [146], the axially wound individual ®bers, the axially wound fabric, and the rectangular modules. The four correlations were indistinguishable within experimental error, so the following average correlation was proposed: Sh ˆ 0:82Re0:49 Sc0:33 :

(36)

Wang and Cussler [137] also compared gas stripping with hollow ®ber contactors containing a woven fabric of microporous polypropylene ®bers to

81

Fig. 9. A rectangular module containing two baffles (redrawn from [137]).

modules built from individual ®bers. The three module con®gurations they used are shown in Figs. 2, 9 and Fig. 10. The ®rst (Fig. 2) was a conventional parallel ¯ow module containing individual ®bers; liquid and gas were placed on the tube and shell sides, respectively. The second and third con®gurations (Figs. 9 and 10) used the woven fabric instead of individual ®bers, with shell side ¯uid ¯ow primarily across rather than parallel to the ®bers. The module depicted in Fig. 9 was rectangular with two baf¯es; liquid and gas ¯owed through the shell and ®bers, respectively. The third con®guration (Fig. 10) was a fully baf¯ed cylindrical module; again liquid and gas were placed on the shell and tube sides. By performing a mass balance around a differential module segment, the authors showed that the mass transfer coef®cient for removal of oxygen from water

Fig. 10. A fully baffled cylindrical module (redrawn from [137]).

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A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

into nitrogen can be described by k ˆ ÿQ=A ln…C1 =C0 †:

(37)

Here Q is the liquid ¯ow rate, A is the membrane area, and C1 and C0 are the liquid outlet and inlet concentrations, respectively. The authors measured C1 and C0, and then used this equation to determine the mass transfer coef®cient for a given ¯ow rate and module geometry. The Sherwood number correlation obtained for the rectangular module (Eq. (38)) was similar to the one reported by Wickramasinghe et al. [140] for commercial oxygenators (Eq. (28a)): Sh ˆ 0:18Re0:86 Sc0:33 :

(38)

Since commercial oxygenators are often compromised by uneven ¯ow as discussed above, the authors surmised that the problem also existed in their rectangular module. Mass transfer coef®cients for the cylindrical modules were higher, possibly because the rectangular module contained stagnant liquid zones between adjacent ®bers in each layer of fabric. Coef®cients for the cylindrical modules were closer to the results obtained by others using either hand-built, carefully made axially wound, or other cylindrically wound baf¯ed fabric module designs [141,146]. This indicates that the authors were successful in achieving uniform ¯ow in the cylindrical modules by using the woven fabric, and the proposed Sherwood number correlation is in fact similar to other correlations obtained with modules having good ®ber spacing (e.g., Eq. (36)): Sh ˆ 0:46Re0:46 Sc0:33 :

(39)

Bhaumik et al. [13] measured mass transfer coef®cients obtained upon absorption of CO2 from pure CO2 and from a CO2/N2 mixture into water using a fabric made from microporous polypropylene hollow ®bers. The fabric was wound helically around a central core similar to the con®guration depicted in Fig. 6. Gas ¯owed on the tube side, while liquid entering the central core was forced radially outward (perpendicular to the ®bers) by the plug. The number of ®ber wraps around the central core was kept small so that the liquid super®cial velocity was clearly de®ned. This in turn was expected to give the correct mass transfer correlation rather than the module average value one

would expect with many wraps. Modules with two or seven wraps were studied using Reynolds numbers ranging from 0.02 to 1.01. The similarity of data obtained with pure CO2 and mixed gas plus the observed effect of changes in liquid ¯ow rate led the authors to conclude that gas phase resistances were negligible. Data obtained from both modules using pure CO2 or mixed gas were ®tted reasonably well by the following correlation: Sh ˆ 0:57Re0:31 Sc0:33 :

(40)

The authors compared their data to the predictions of other correlations developed for low Reynolds numbers. There was a good ®t with the correlation of Wickramasinghe et al. [141], who stripped oxygen from water using a similar module and ¯ow pattern (Eq. (36)). 5.3. Mass transfer with chemical reaction As in any mass transfer device, chemical reaction can have a major effect on the rate of mass transfer in membrane contactors. Consider gas/liquid mass transfer: in the absence of chemical reaction, the gas phase resistances are usually negligible and the liquid boundary layer resistance controls. However, if a chemical reaction occurs between the absorbed species and the solvent, the liquid side resistance is reduced and the gas phase resistance may actually control. Suppose we have the following reaction: A …from gas† ‡ bB …liquid† ! products …liquid† (41) For a fast reaction and low solute concentration in the liquid, CB is essentially constant and the mass transfer coef®cient is given by k ˆ …kR DAL †1=2 ;

(42)

where kR is the ®rst order reaction rate constant. Here the reaction is fast but not so fast that the reactants cannot coexist; an example is the mass transfer of chlorine into water, where the two react to form hypochlorous and hydrochloric acids [70,102]. For an instantaneous reaction with negligible gas phase resistance, k0 ˆ k…1 ‡ DBL CB =bDAL CAi †;

(43)

A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

where k0 and k are the mass transfer coef®cients with and without chemical reaction, respectively, and i refers to the ¯uid/¯uid interface. Here the reaction is so fast that the reactants cannot coexist, e.g., the absorption of hydrogen sul®de in aqueous sodium hydroxide to produce sodium bisul®de [70,102]. Comparable equations can be written for liquid/liquid mass transfer with reaction. Reports dealing with chemical reactions in liquid/ liquid systems include Refs. [8,10,50,72,78,104,148± 151]. Gas/liquid reacting systems have been discussed by Qi and Cussler [97,98], Cooney and Jackson [21], Semmens et al. [112], Karoor and Sirkar [56], and Kreulen et al. [65]. Examples of reacting systems are given in Section 6. 5.4. Additional considerations in contactor design and operation 5.4.1. Design objective In mass transfer with membrane contactors the design objective depends on the application. For example, with blood oxygenators the objective is to maximize the amount of oxygen transferred per unit volume. This minimizes the contactor volume required, which in turn minimizes the volume of blood outside the body. On the other hand, for industrial applications the objective is clearly to minimize the cost per amount of mass transferred, which does not necessarily coincide with the maximum solute transferred per unit volume [139]. 5.4.2. Design approach Module length is a key design speci®cation because it is directly related to the number of transfer units (see Eq. (7)). However, the higher ef®ciency offered by longer modules comes at the cost of increased pressure drop. If the module is too long, the pressure of the nonwetting ¯uid may fall below that of the wetting ¯uid, allowing bulk ¯ow of the wetting into the non-wetting ¯uid (see Section 3.2). There may be some recourse in increasing the inlet pressure, but that pressure is limited by the breakthrough pressure. A number of design features must be considered to obtain the required capacity and number of transfer units while meeting these pressure constraints, i.e., tube diameter, wall thickness, porosity, tortuosity, packing factor, ¯ow rates, inlet concentrations, dis-

83

tribution coef®cient, and ¯uid physical properties. Prasad and Sirkar [96] suggested a design procedure that considers these factors and gives the required number of ®bers, module diameter and length for a given design problem. For cases where single module designs are not practical, Prasad and Sirkar suggested the use of series/parallel cascades. Reed et al. [102] gave several design examples. 5.4.3. Optimum fiber diameter Wickramasinghe et al. [139] looked at the cost of stripping gas from water into nitrogen using a hollow ®ber contactor with the water ¯owing through the tubes. Both a highly volatile and a less volatile solute were studied; with oxygen, the volatile solute, the tube side boundary layer resistance controlled, while with bromoform, the less volatile solute, the shell side resistance was important. Cost per unit mass transferred was plotted vs. ®ber diameter, with either membrane cost per unit area (for a ®xed module length) or module length (for a ®xed membrane cost) as a parameter. The cost was taken as the sum of the membrane and pumping costs; the amount of mass transferred was calculated using shell and tube side mass transfer coef®cients obtained from Eqs. (20) and (16), respectively. Results are shown in Figs. 11±14. In nearly all cases there was an optimum ®ber diameter which corresponded to a minimum cost. That is, pumping cost decreased with increasing ®ber diameter, but membrane cost increased due to the increase in membrane area. This means that pumping costs dominated at small diameters while membrane costs dominated at large ones, hence the existence of an optimum. The cost varied as dÿ4 at low diameters and linearly with d at large diameters. Not surprisingly, an increase in membrane cost decreased the optimum ®ber diameter. Furthermore, the optimum diameter increased with increasing length; this was explained by the increase in pumping cost resulting from the increase in length, which in turn required a larger diameter before the membrane cost began to dominate. It is interesting that in most cases the optimum diameter was around a few hundred microns, which coincides with the diameters of commercially available ®bers. While the trends were the same with oxygen and bromoform, the costs were about half for bromoform under comparable conditions. This resulted from the

84

A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

Fig. 11. Cost per unit mass transferred for stripping oxygen from water into nitrogen using a module 1 m in length. The parameter is membrane cost in $/m2 year (redrawn from [139]).

Fig. 13. Cost per unit mass transferred for bromoform stripping. Again the module is 1 m in length, and the parameter is membrane cost in $/m2 year (redrawn from [139]).

difference in controlling resistance. That is, since the tube side resistance controlled for the oxygen but not for the bromoform, a higher mass transfer coef®cient could be achieved with the latter by increasing the

shell side ¯ow rate. The increase in mass transfer rate more than compensated for the increase in pumping cost corresponding to the higher ¯ow rate, hence the decrease in cost per unit mass transferred. However,

Fig. 12. Cost per unit mass transferred for oxygen stripping. This figure shows the effect of module length for a fixed membrane cost of $10/m2 year (redrawn from [139]).

Fig. 14. Cost per unit mass transferred for bromoform stripping. This figure shows the effect of module length for a fixed membrane cost of $10/m2 year (redrawn from [139]).

A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

the calculations for bromoform were the least reliable because shell side ¯ow is not as well understood as the tube side ¯ow and the available correlations for the mass transfer coef®cient are not as good, as discussed above. 5.4.4. Flow direction Wang and Cussler [137] explained that conventional parallel ¯ow modules offer true countercurrent ¯ow and are preferred when the membrane or the tube side boundary layer resistance controls. However, with these modules mass transfer coef®cients can be reduced and/or ¯ows can become uneven if the shell side resistance is signi®cant; in this case a cross¯ow design is preferred. Flow normal rather than parallel to the ®bers leads to higher mass transfer coef®cients, but the price is loss of ef®ciency compared to countercurrent designs. Some ef®ciency is regained by the use of baf¯ed modules, which provide elements of both countercurrent and cross¯ow; ef®ciency increases as the number of baf¯es increases, but pressure drop increases as well, and the modules become more dif®cult to build. Wang and Cussler [137] looked at the effect of the number of baf¯es on mass transfer performance. They measured the fraction of toluene stripped from water into nitrogen using a rectangular module (Fig. 9) equipped with one, two or four baf¯es, operated in either cocurrent or countercurrent mode with the liquid on the tube side. Toluene removal was higher for countercurrent ¯ow as expected. Interestingly, oscillations in the fraction of toluene removed vs. number of baf¯es were observed for cocurrent but not countercurrent ¯ow. The authors explained that in cocurrent ¯ow, exit liquid contacts less saturated gas if there is an even number of baf¯es and more saturated gas if there is an odd number. Consequently, more toluene is removed with an even number and less with an odd number, hence the oscillations. No oscillations were observed with countercurrent ¯ow since exiting liquid always contacts the entering gas. The effect of baf¯ing on oxygen stripping from water into nitrogen using cylindrical modules was assessed with a predictive model. Performance was examined as a function of the absorption factor GH/Q for several different con®gurations, including cross¯ow, countercurrent and cocurrent ¯ow, all with the gas ¯ow inside the ®bers. The latter two contained

85

zero, two or ®ve baf¯es; no baf¯es were used in the cross¯ow case. The highest oxygen removal was achieved using countercurrent ¯ow with ®ve baf¯es, although performance with two baf¯es was almost as good. Cross¯ow with no baf¯es was superior to cocurrent ¯ow with two. 6. Applications A summary of applications is presented in Table 3. 6.1. Liquid/liquid extraction The versatility of membrane contactors is clear from the abundance of published reports on liquid/ liquid extraction applications, and several of these have already been cited. Results from two comprehensive studies [91,93] are summarized in Table 4. These works illustrate the utility of the technology for chemical systems covering a broad range of mi values using different membrane geometries, materials and ¯ow con®gurations. In liquid/liquid extraction using hollow ®ber modules, several considerations govern which ¯uid is placed on the shell side and which one is selected for ¯ow through the ®bers. Because the shell side ¯uid is subject to bypassing and channeling as discussed above, a high degree of extraction is often dif®cult to obtain with the solute-containing feed on the shell side. However, if the feed contains particulates that are not at least one order of magnitude smaller than the inside diameter of the ®bers, then tube side ¯ow is not advisable. On the other hand, large particulates are more likely to settle out in dead spots on the shell side, which suggests particle size reduction (e.g., with a homogenizer or grinder) or removal by ®ltration to allow tube side ¯ow. Another consideration is the integrity of the bond between the ®bers and the tubesheet; if leakage is a concern, the lower pressure ¯uid (i.e., the one which wets the pores) preferably ¯ows on the tube side [22,96]. In most liquid/liquid extraction applications the lowest membrane mass transfer resistance is obtained if the pores are ®lled with the ¯uid in which the solute is most soluble. This suggests the use of a hydrophilic or hydrophobic membrane for a low or high mi solute, respectively (see Eqs. (2)±(5)). That is, a low mi solute

86

Table 3 Summary of applications Comments

Selected references

Liquid/liquid extraction

Membrane contactors have been employed in a variety of chemical systems covering a broad range of mi values. A number of membrane geometries, materials of construction and flow configurations have been investigated. Results from two comprehensive studies are given in Table 4 A variety of solute/solvent systems have been studied, some including chemical reactions. Solutes investigated in recent work include CO2, SO2, NOx, Br2, H2S, I2, N2, NH3, (NH4)2S, acetic acid, HCl, lactic acid, ethane, ethylene and VOCs. Commercial applications include carbonation of beverages, treatment of boiler feedwater, stripping of CO2 from anion exchange feed streams (to extend bed life), nitrogenation of beer to provide a dense foam head, and deoxygenation of beer to preserve flavor. Wastewater treatment and semiconductor applications are discussed separately Membrane contactors offer a low-cost alternative for production of otherwise expensive supercritical extracts, particularly those of interest to the food and flavor industry Selective extraction of d-leucine from a racemic mixture in water using a 1-octanol solution of N-n-dodecyl-l-hydroxyproline has been demonstrated. A membrane reactor has been employed to selectively remove the undesirable stereoisomer from a racemic mixture of enantiomers encountered in the production of a drug used to treat hypertension and angina. Several investigators have shown that productivity of ethanol and other metabolic products can be increased by in situ extraction using membrane contactors. Enzymatic hydrolysis of triglycerides with simultaneous separation of the resulting glycerol and fatty acids has also been demonstrated Various proteins have been investigated, including cytochrome-c, myoglobin, a-chymotrypsin, catalase, urease, and a-amylase Several important compounds (e.g., 4-methylthiazole, 4-cyanothiazole and mevinolinic acid) have been extracted from actual industrial steams using membrane contactors Recovery of a number of troublesome pollutants from aqueous streams has been demonstrated, including 2-chlorophenol, benzene, nitrobenzene, trichloromethane, tetrachloromethane and acrylonitrile. Bubble-free aeration offers several advantages over conventional approaches, including no venting of VOCs and 100% oxygen transfer efficiency at low power input The efficacy of the technology has been demonstrated for recovery of a variety of ions from aqueous wastewaters. Some authors have reported the use of dual contactors operated in chelating and stripping mode, respectively. Numerous applications using chemical reaction to facilitate ion removal have been described Membrane contactors are used in semiconductor manufacturing operations for the production of ultrapure water and for ozonation of cleaning water In this process, one or more volatile components of a liquid feed are transferred through a non-wetted microporous membrane into another liquid. Applications include fruit and vegetable juice concentration and production of low alcohol wine

[8,10,22,34,59,91,93,96,101,109]

Gas absorption and stripping

Dense gas extraction Chiral separations

Fermentation and enzymatic transformation Protein extraction Pharmaceutical applications Wastewater treatment

Metal ion extraction

Semiconductors Osmotic distillation

[12,52,56,64,65,88,97±100,118, 122,125,134,140,146]

[103,121] [35,72]

[19,40,41,50,54,55,74,78,79, 119,136] [29,33] [94,95] [1,2,25,26,75,83,91,93,96, 111,113±115,152,153] [20,57,96,148,150,151]

[23,31,45,143] [37,49,68,87]

A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

Application

A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

87

Table 4 Liquid/liquid extraction using membrane contactorsa Feed solvent n-Heptane Water Water Water Water Water Water Water Water Watere Watere

Solute Toluene Acetic acid Acetic acid Acetic acid Acetic acid Acetic acid Acetic acid Succinic acid Succinic acid Phenol Phenol

Extraction solvent b

NMP MIBK MIBK MIBK MIBK Xylene Xylene n-Butanol n-Butanol MIBK MIBK

Partition coefficient 1.03 0.52 0.52 0.52 0.52 0.0125d 0.0125d 1.49 1.49 119 119

Membrane configuration

Membrane material

Flat Flat Flat Hollow Hollow Hollow Hollow Hollow Hollow Hollow Hollow

Hydrophobic Hydrophilic Compositec Hydrophobic Hydrophilic Hydrophobic Hydrophilic Hydrophobic Hydrophilic Hydrophobic Hydrophilic

fiber fiber fiber fiber fiber fiber fiber fiber

a

Refs. [91,93]. A polar organic solvent. c A hydrophobic sheet placed on the top of a hydrophilic sheet. d Prasad and Sirkar [93] gave a range of 0.0125±0.039 corresponding to aqueous concentrations of 0.27±4.75 mol/cm3. e pH 6.8. b

is more soluble in a polar solvent than in a non-polar one, and the polar solvent will wet the pores of a hydrophilic membrane; the situation is analogous for a high mi solute and a hydrophobic membrane. However, with proper membrane pretreatment, a polar solvent will wet a hydrophobic membrane, so that low membrane resistance can still be achieved with a low mi solute; similarly, hydrophilic membrane pores can be ®lled with a non-polar solvent [90]. For example, Prasad et al. [90] obtained a ¯at Cuprophan 150PM membrane (regenerated cellulose, ENKA) from the manufacturer with the pores ®lled with a low-volatility organic; the wetting characteristics of the membrane were set by the method used to remove this organic. That is, operation with water®lled pores was achieved by ®rst washing the membrane with isopropyl alcohol then air-drying; conversely, operation with a non-polar organic in the pores was accomplished by ®rst washing with xylene. Similar results were obtained with other hydrophilic membranes including nylon (polyamide, ¯at, ENKA), ceramic (alumina, tubular, ALCOA), and glass (tubular, ASAHI). A hydrophobic membrane (Celgard 2400, polypropylene, tubular, CELGARD LLC) was also operated with either a non-polar (n-heptane) or polar (n-methyl pyrrolidone) solvent in the pores by soaking in the respective solvent before use. This unconventional approach allows one to obtain the lowest possible mass transfer resistance with

fewer restrictions on membrane material. However, a penalty is paid in breakthrough pressure, which is signi®cantly lower for hydrophilic and hydrophobic membranes ®lled with non-polar and polar ¯uids, respectively [90]. Reed et al. [101] performed pilot scale extractions of chlorinated and aromatic compounds from industrial wastewaters in Holland. Some 75 l/h of wastewater were treated using three Liqui-Cel modules (Fig. 1), each with an interfacial area of 2±3 m2. Treatment was performed successfully for periods up to three months, with individual contaminant levels reduced to less than 10 mg/l, and process economics were competitive with other approaches. Basu and Sirkar [8] looked at liquid/liquid extraction facilitated by chemical reaction. These investigators used a hydrophobic ¯at sheet (i.e., organic®lled pores) to study the extraction of phenol from MIBK into aqueous sodium hydroxide, which is facilitated by the following instantaneous reaction: C6 H5 OH ‡ NaOH ! C6 H5 ONa ‡ H2 O

(44)

For sodium hydroxide concentrations below some critical value, the reaction front resided within the aqueous boundary layer; the front moved closer to the liquid/liquid interface as the NaOH concentration increased, ®nally reaching the interface when some critical concentration was exceeded. Basu and Sirkar presented the entire set of resistance-in-series equa-

88

A. Gabelman, S.-T. Hwang / Journal of Membrane Science 159 (1999) 61±106

tions for mass transfer with instantaneous reaction in the aqueous boundary layer. Cases considered involved hydrophobic, hydrophilic and composite membranes, each with aqueous reactant concentrations above and below the critical value. In a related study, Basu and Sirkar [10] used hydrophobic hollow ®ber modules to investigate the extraction of citric acid from water into trioctyl amine/ MIBK followed by back-extraction into sodium hydroxide. The formation of the citric acid/trioctyl amine complex followed the fast reversible reaction C6 H8 O7 …aq† ‡ 3R3 N …org† $ …R3 NH†3 C6 H5 O7 …org†

(45)

The effects of aqueous phase ¯ow rate, organic phase ¯ow rate and trioctyl amine concentration were studied. Experiments were performed with the aqueous phase ¯ow on the tube side as well as the shell side, and a mathematical model was proposed that successfully predicted the effect of the chemical reaction on the mass transfer rate. Under most conditions, the aqueous boundary layer controlled as expected for a hydrophobic membrane and a high mi solute (see Eq. (3)), i.e., the citric acid/amine complex. However, in some cases, the size of the complex rendered the membrane resistance signi®cant. Interestingly, the back-extraction into NaOH was much slower than the forward extraction into the organic phase. Keurentjes et al. [59] employed a membrane contactor to extract a series of fatty acids from soybean oil into 1,2-butanediol using cellulosic hollow ®bers with the oil ¯owing on the tube side. The fatty acids were recovered by combining the extract with water, separating the resulting dispersion into pure fatty acids and a fatty acid-free 1,2-butanediol/water mixture, then dewatering the mixture to recover 1,2-butanediol for reuse. The chain length of the fatty acids studied ranged from 6 (caproic) to 22 (erucic). In most cases the major resistance to mass transfer was diffusion through the membrane pores, which were wetted by the 1,2-butanediol. Because of the low interfacial tension with this system, small pores were necessary to avoid breakthrough (see Eq. (6a)). The authors found that the mass transfer coef®cient increased and the equilibrium distribution shifted more towards the 1,2-butanediol with decreasing chain length. That is, there were two mechanisms

acting in the same direction to obtain enhanced selectivity, which suggests the use of a membrane contactor for fractionation of fatty acids. The authors explained that higher purities can be obtained in less time with this approach than with traditional crystallization methods. Two other examples of liquid/liquid extraction applications were cited by Sirkar [125]:  Matsumura [73] used a hydrophilic membrane to study the extraction of ethanol, isopropanol and nbutanol.  Brose [15] extracted onion oil and other flavors from fruit and vegetable processing waste streams. Small membrane contactors have been tested by a number of industrial organizations for a variety of liquid/liquid extraction applications. However, large systems are not yet ready for use, primarily because of a lack of suitable materials of construction [125,126]. 6.2. Gas absorption and stripping Even though the large density difference between liquid and gas leads to faster countercurrent ¯ows and somewhat alleviates the problems of ¯ooding and unloading in conventional equipment, the advantages of membrane contactors cited above are still important [28]. Indeed, gas/liquid contacting applications are enjoying numerous commercial successes; several of these are discussed below. Two modes of operation are possible in gas/liquid contactors: wetted mode and dry (or non-wetted) mode. Wetted mode occurs when the pores are ®lled with liquid, e.g., if the liquid phase is aqueous and a hydrophilic membrane is used. Conversely, a hydrophobic membrane would operate in dry mode in this case because the pores would be ®lled with gas. Dry mode is usually preferred to take advantage of the higher diffusivity in the gas; however, wetted mode may be preferred if there is a fast or instantaneous liquid phase reaction and as a result the gas phase resistance controls [56]. Karoor and Sirkar [56] studied the absorption of pure CO2, pure SO2, CO2 from CO2/N2 mixtures, and SO2 from SO2/air mixtures into water. They used a parallel ¯ow module employing microporous polypropylene ®bers, usually with the water on the tube side. These researchers also modeled the system by writing the differential mass balance, substituting the

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appropriate velocity pro®les and reaction kinetics, then solving using a ®nite difference scheme [30]. A parabolic velocity pro®le was assumed for the tube side, while the shell side velocity pro®le was predicted using Happel's free surface model [48]. Both wetted and dry modes were investigated. Absorption of SO2 was facilitated by the following chemical reaction: SO2 ‡ 2H2 O $ H3 O‡ ‡ HSOÿ 3;

KE ˆ 0:014 M (46)

Further disassociation of HSOÿ 3 was avoided by proper selection of pH. The reaction of CO2 with water was neglected because of the low value of the equilibrium constant. As expected, CO2 absorption was limited by the liquid phase resistance, and best results were obtained in the dry mode. Experimentally determined mass transfer coef®cients agreed well with those predicted by the model when the aqueous phase ¯owed through the tubes, and the LeÂveÃque [71] solution was also in good agreement with the data. The mass transfer coef®cients were as much as ®ve times those typically obtained for conventional packed towers. Mass transfer coef®cients obtained with SO2 absorption also agreed well with predicted values for liquid ¯ow through the tubes, and those values were 10 times higher than those usually obtained in packed towers. In spite of the chemical reaction (Eq. (46)), resistance was somewhat higher in the wetted vs. dry operating mode with SO2 absorption. On the other hand, Yang and Cussler [146] did in fact show a shift in controlling resistance that was attributable to chemical reaction. That is, when oxygen on the shell side of a microporous polypropylene hollow ®ber module was absorbed into water on the tube side while operating in dry mode, the aqueous phase boundary layer resistance controlled as expected. However, with CO2 absorption into dilute sodium hydroxide, the chemical reaction reduced the tube side mass transfer resistance so that it no longer controlled. Qi and Cussler [97,98] also reported on absorption of CO2 into aqueous sodium hydroxide using hydrophobic hollow ®bers, including characterization of the mass transfer behavior. Jansen et al. [52] employed microporous hydrophobic hollow ®bers (diameter: 0.6 mm; length: 25 cm) to absorb SO2 from (a) a model gas of SO2 in nitrogen, and (b) real ¯ue gas from a coal-®red

89

boiler; the reactive absorbent was Na2SO3. Over 99% recovery was obtained with the model gas, and the results agreed well with theoretical predictions. Similar recoveries were maintained with the real ¯ue gas over a period of 500 h, and there were no problems with dust, particles, condensation, or other ¯ue gas components (e.g., NOx, CO2, HCl). Encouraged by this success, a pilot plant was built in Krim, Holland, as a collaborative effort between the TNO Institute of Environmental and Energy Technology, Hoogovens Technical Services, Bronswerk Heat Transfer and AVEBE. The plant was tested for two six-month seasons; SO2 recovery exceeded 95%, there were no fouling problems and in general the plant was easy to operate. Bhaumik et al. [12] investigated the use of a microporous polypropylene membrane contactor operated with a pressure swing cycle, a technique known as rapid pressure swing absorption, to recover CO2 from nitrogen. In this work the absorbent (20% aqueous diethanolamine, or DEA) was stationary on the shell side. Feed gas (10% CO2 in nitrogen) was passed through the ®bers, CO2 was absorbed and the puri®ed gas was obtained at the ®ber outlet. At the end of the absorption period (10±25 s), absorbed gas was removed from the absorbent by pulling a vacuum on the shell side for 15±20 s, then the cycle was repeated. This technique offers a number of advantages over conventional pressure swing absorption in addition to those already cited for membrane contactors in general:  Purification ability is extraordinary. The exit gas contained less than 50 ppmv (the detection limit for the analytical method used) at a flow rate of 50±55 cm3/min, or 1.1±1.2 cm3/min per cm3 of module volume.  The capacity and selectivity of liquid absorbents, especially reactive ones, are much higher than for most solid adsorbents.  The absorber and stripper used in conventional processes are replaced with one vessel.  High rates of absorption/desorption are possible by using short cycle times. The authors pointed out that the same approach is applicable to a variety of gas puri®cation problems, including removal of CO2 from ¯ue gas, CO and CO2 from post-shift reactor synthesis gas, H2S from streams containing H2S and CO2, and VOCs from air.

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Qi and Cussler [99,100] studied the transport of nine volatile solutes from aqueous solution on the tube side through a hydrophobic gas-®lled membrane and into a second aqueous solution on the shell side. In this manner stripping and absorption were carried out in a single device. The nine solutes were Br2, H2S, I2, SO2, NH3, (NH4)2S, acetic acid, HCl, and lactic acid; mass transfer coef®cients ranged from about 0 for HCl and lactic acid to 0.00041 cm/s for Br2. Although the ¯ux depended both on the volatility of the solute and its solubility in water, the latter was more important. In some cases the absorption step was facilitated by reaction with the stripping solution, e.g., 0.5 N H2SO4 was used to strip NH3; the relative resistance, of the shell side boundary layer diminished accordingly. The authors explained that ionization of solute can have a dramatic effect on mass transfer resistance so that the overall resistance is no longer simply the sum of the boundary layer and membrane resistances. Kreulen et al. [64] studied the absorption of CO2 into water/glycerol using polypropylene or polysulfone hollow ®bers. Interestingly, when the liquid ¯owed through the ®bers, agreement between theoretical and experimental values of the mass transfer coef®cients was obtained when the active mass transfer area was taken as the entire area of the ®bers, not just the area of the pores themselves. This was true regardless of the porosity of the ®bers, which was 0.7 and <0.03 for the polypropylene and polysulfone, respectively (pore sizes were 110ÿ7 and 510ÿ9 m). The authors explained that the distance between adjacent pores was much less than the distance from the ®ber wall to the center of the ®ber. This meant that diffusion parallel to the ®ber led to a boundary layer that was in effect homogeneous and saturated, hence the importance of the entire membrane area rather than just the pore area. Other ®ndings from this work were as follows:  For liquid flow through the tubes, mass transfer coefficients followed the Graetz±LeÂveÃque solution for heat transfer [46,47,71] (see Eq. (16)).  For liquid flowing around regularly packed fibers, mass transfer was described reasonably well by the correlation of Miyatake and Iwashita [77], which was derived for a similar heat transfer problem.  Correlating mass transfer results obtained with liquid flowing around irregularly packed fibers

was difficult because of the unknown dimensions of the channels between the fibers. Yang and Cussler [147] demonstrated the ef®cacy of two types of microporous polypropylene membrane modules (parallel ¯ow and cross¯ow) as arti®cial gills. Water was pumped through the tubes of the module, while gas was pumped in a closed loop between the module shell and a small box. Transfer of oxygen from the water to the gas and CO2 from the gas to the water was actually suf®cient to sustain a small animal placed in the box. When a hamster was put in the air-®lled box, the oxygen and CO2 concentrations in the box reached a steady state (corresponding to the hamster's metabolic rate) in about 2 h. A white rat remained healthy in the box for three days. To accommodate the larger size and higher metabolic rate of a dog, a larger box was used and 40 modules were connected in parallel; the dog, a 4.9 kg West Highland terrier, survived quite nicely. The arti®cial gill was even demonstrated on a 57 kg human (the ®rst author of the paper); for that experiment the pump and box were replaced by a SCUBA mouthpiece, hoses and a rubber bladder. The authors presented equations that successfully predicted the oxygen and CO2 concentrations in the box over time as well as the mass transfer coef®cients. They concluded that arti®cial gills could be designed which would keep small animals alive. Such gills could also keep resting humans alive if water were pumped rapidly through the gill, but the task would be more challenging if the human moved through the water. Jansen et al. [52] proposed the use of membrane contactors to remove harmful components of environmental tobacco smoke (ETS) from the air, e.g., in a home or of®ce where smokers are present. These investigators used a smoke generator to produce ETS, then passed the contaminated air through a laboratory membrane module with water ¯owing on the opposite side. Percent removal ef®ciencies were as follows: acetone 97, styrene 15, formaldehyde 98, nicotine 99, and ammonia 95. The poor recovery for styrene is not surprising considering its poor solubility in water. Tsou et al. [134] investigated the use of membrane contactors for separation of ole®ns from petrochemical feedstocks, using silver as a complexing agent to improve the ¯ux and selectivity [43,106]. They fed a

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mixture of 74/26 ethylene/ethane to the shell side of a hydrophilic hollow ®ber membrane module with a silver nitrate solution (0.5±5 N) on the tube side. Shell and tube side pressures of 0±200 psig were possible using back pressure regulators. The exit liquid stream ¯owed to a ¯ash pot aided by a helium sweep to release the absorbed gas, and the regenerated absorbent was returned to the module. Throughput and separation ef®ciency were high enough to support commercial implementation of the technology not only to separate ethylene from ethane but for puri®cation of other ole®ns as well. The authors have used the technology to reliably produce puri®ed propylene from re®nery grade propylene at pilot scale, i.e., one half barrel per day. Poddar et al. [88] looked at absorption of volatile organic compounds (VOCs) from N2/air into silicone oil or mineral oil. Two types of hollow ®ber modules were used: one containing a porous hydrophobic membrane, and one containing a highly VOC-permeable non-porous silicone skin over a porous membrane; in both the cases the absorbent ¯owed through the shell side. Spent absorbent was regenerated under vacuum in a separate module then recycled. Ef®cient VOC removal from the gas stream was achieved; the highest ef®ciency was obtained with toluene, followed by methylene chloride, acetone and methanol. Ef®ciency was higher with silicone vs. mineral oil because of the higher diffusivities and mi values with the former, but mineral oil provided better long term stability. The absorbent-®lled porous membrane was responsible for 80±95% of the mass transfer resistance, while for the skinned membrane most of the resistance was found in the skin and the liquid boundary layer outside of it. Plots of dimensionless concentration vs. the inverse of the tube side Graetz number agreed well with theoretical predictions. Mass transfer coef®cients were substantially higher than those typically obtained with packed towers. The authors suggested that their regenerative membrane-based absorption process could eventually replace carbon adsorption in VOC removal applications. Similarly, Jansen et al. [52] discussed the use of a hydrophobic membrane coated with cross-linked polydimethylsiloxane (PDMS) for VOC recovery from air into non-polar hydrocarbon or silicone oil. This approach (known as selective membrane absorp-

91

tion) has been successfully demonstrated with the feed gas having a VOC content of only 0.6 mg/m3. Gas absorption/stripping with membrane contactors was recently reviewed by Sirkar [122]; additional information can be found there. Gas/liquid contacting in wastewater treatment and semiconductor manufacturing are discussed in Sections 6.8 and 6.10, respectively. 6.2.1. Commercial applications A number of gas/liquid applications have been commercialized successfully. For example, a Pepsi bottling plant in West Virginia has operated a bubblefree membrane-based carbonation line since December 1993. The plant processes 112 gal/min of beverage with 10 4 in.28 in. Liqui-Cel Extra Flow modules (see Section 2 and Fig. 1) having a total interfacial area of 193 m2. Advantages over conventional carbonation methods include reduced foaming, improved yield, lower CO2 pressures, and minimal drop in ®ller speed when ®lling at elevated temperatures [51,86,125]. Liqui-Cel contactors are also used in beer production. Several plants are using the technology for CO2 removal followed by nitrogenation to obtain a dense foam head, and others are employing membranes to remove oxygen from beer to preserve ¯avor. Some breweries are using Liqui-Cel equipment to strip oxygen from water, then using the deoxygenated water to dilute beer that contains 9±10% alcohol. The high alcohol content is obtained from high gravity brewing, a process that offers the brewer increased capacity [86]. Other commercial applications include treatment of boiler feedwater, stripping of CO2 from anion exchange feed streams to extend bed life, and production of ultrapure water for semiconductor manufacturing (see Section 6.10) [5,118,126]. Sengupta et al. [118] reviewed three operating modes for stripping of volatile gases from water for these applications: vacuum, inert sweep or purge gas, and a combination of the two; they also discussed performance data, mass transfer correlations and cost/performance characteristics. 6.3. Dense gas extraction Like liquids, dense gases offer a high solubility of many solutes of interest, yet they also offer the high

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mass transfer rate and low pressure drop enjoyed with gases. Furthermore, solubility (or distribution coef®cient) is usually a strong function of density so that the dense gas is easily separated from the solutes simply by reducing the pressure [14,129]. This is particularly attractive in food and ¯avor applications where solvent residues (e.g., in the production of spice oleoresins using hexane extraction) are coming under increased scrutiny by the FDA and other regulatory organizations. Dense CO2 is a particularly good choice for applications like these [80], not only because it leaves no residue but also because it is non-toxic, non-¯ammable and inexpensive. The lack of toxicity offers an important environmental advantage in today's stringent regulatory climate as well as increased ease of handling. The strong dependence of solubility on density allows one to selectively extract solutes of interest or to fractionate them. For example, one may be interested in the ¯avor but not the fat contained in an essential oil such as peanut oil for use in a low-fat food application. In this case a CO2 density can be selected (via the temperature and pressure) so that the ¯avor is largely extracted yet the fat is largely excluded, since the latter is less soluble at a given density. Alternatively, both the fat and ¯avor fractions may be of interest but for different applications. In this case one would operate at a high density so that both the fat and ¯avor are extracted, then collect separate fat and ¯avor fractions by letting the pressure down in stages (i.e., collect the fat at an intermediate pressure and the ¯avor at a low pressure). Although CO2 is often the solvent of choice in dense gas extraction, other solvents are also used. For example, in many cases propane offers a more complete extraction than one would obtain with CO2. CF Foods (Arvada, CO) has developed a process to remove nearly 100% of the oil from food products like peanut and cocoa using dense propane, obtaining both ¯our and oil as valuable products [85]. Of course, the ¯ammability of propane is a serious disadvantage. More generally, however, the disadvantage of handling CO2 and other ¯uids in the near critical or supercritical range is the high pressure required. For example, the critical pressures for CO2 and propane are 73.8 and 42.5 bar, respectively, as given by Smith and Van Ness [128] and other textbooks on thermodynamics. Operating at these pressures is expensive.

That is, capital cost is high because the equipment must be designed to withstand high pressure; operating cost is also high because the capital cost leads to a sizeable depreciation charge, and the compression of the gas incurs substantial charges for power and maintenance. Fortunately, the issue of cost can be addressed by using a membrane contactor for dense gas extraction, a patented approach given the name PorocritTM by the inventors [103,121]. A key aspect of the invention is the ability to operate at high pressure using hollow ®bers having a burst strength of <100 psig. This is accomplished by a novel design that equalizes the pressure on either side of the membrane so that even though the system pressure is quite high, the pressure drop across the membrane is nearly zero. The inventors pointed out that the lower viscosity of CO2 results in better shell side distribution and less bypassing than one would obtain with more traditional liquid solvents. The advantages of membrane contactors cited above, particularly the high ef®ciency (low HTU) afforded by the plentiful interfacial area, allow a given supercritical extraction task to be performed in considerably smaller (hence less expensive) equipment than would be required using a traditional contacting column. Apparently, little work has been reported on determination of mass transfer coef®cients for liquid/dense gas systems in membrane contactors. Brunner [16] cited the work of Zehnder and Trepp [154], who measured mass transfer with a-tocopherol and dense CO2 in laminar ¯ow through tubes, but he stated that in general there were not enough data for development of generally applicable correlations. Such correlations are needed for design if the technology is to realize its full commercial potential; unfortunately, the unique characteristics of dense gases render equations like Eq. (15) unsuitable. One such characteristic is the kinematic viscosity, which tends to be rather low for dense gases. For example, the kinematic viscosity of CO2 at 310 K and 150 bar is one and three orders of magnitude lower than that of water and air, respectively, at 298 K and 1 bar. Since the magnitude of buoyant forces is inversely proportional to the kinematic viscosity, natural convection becomes important near the critical point of a ¯uid [16,32,131]. However, correlations of the type given in Eq. (15) were developed using systems

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where natural convection was not important. To render such correlations valid for dense gases, the Grashof number Gr must be included [11,32]. Gr, the ratio of viscous forces to momentum, arises in similarity solutions to the differential equations that describe mass transfer with natural convection [42]. 6.4. Chiral separations Ding et al. [35] demonstrated the selective extraction of d-leucine from a racemic mixture in water using a 1-octanol solution of N-n-dodecyl-l-hydroxyproline. These investigators were able to obtain nearly complete separation of the two enantiomers, and a plot of the fraction of d-leucine extracted vs. the Graetz number agreed closely with the results predicted by theory [36]. A second example of a chiral separation [72] is given in the following section. 6.5. Fermentation and enzymatic transformation As mentioned earlier, membrane contactors offer a means of in situ extraction of chemical or biochemical reaction products, which can improve productivity by shifting equilibrium to the right. This feature is particularly well suited to fermentation processes because membrane contactors can be operated aseptically. Frank and Sirkar [40,41] studied the fermentation of glucose to ethanol in a membrane contactor. An aqueous solution of glucose and other nutrients was pumped through the shell side, which housed yeast cells immobilized on wooden chips, while dibutyl phthalate was passed through the tubes to extract ethanol as it was formed. Frank and Sirkar [41] found that ethanol productivity increased substantially with increased tube side ¯ow rate, e.g., productivity improved by 28% (from 8.65 to 11.08 g/l h) when solvent ¯ow rate was increased from 0.5 to 0.7 ml/min. These workers also demonstrated the utility of the membrane contactor for gas exchange, i.e., supply of oxygen and removal of carbon dioxide. Several other studies of extractive fermentation have also been reported. These include ethanol fermentation with oleyl alcohol [54,55], sec-octanol [136] or tributyl phosphate [19,74] as the extracting solvent, acetone±butanol±ethanol fermentation with 2-ethyl-1-hexanol as the extractant [119], and lactic

93

acid with tri-(n-octyl)phosphine oxide (TOPO, diluted in kerosene) as the solvent [7,105]. Naser and Fournier [79] presented an economic analysis of ethanol production using extractive fermentation. Other investigators have studied membrane reactors using single enzymes rather than whole cells. For example, Lopez and Matson [72] used a microporous hollow ®ber membrane reactor to selectively remove the undesirable stereoisomer from a racemic mixture of enantiomers encountered in the production of diltiazem. This drug, sold in the US as Cardizem1, is used to treat hypertension and angina. The precursors of interest are the methyl esters of ()-trans-4-methoxy-3-phenylglycidic acid (MMPG); the (2R-3S) form yields the diltiazem stereoisomer with highest activity. The racemic mixture can be puri®ed by using a stereoselective lipase to convert the undesirable (2S3R)-trans enantiomer to (2S-3R)-methoxyphenylglycidic acid and methanol. This unstable acid rapidly decarboxylates to form p-methoxyphenylacetaldehyde, a strong inhibitor of the enzyme, but this decomposition can be arrested by the reversible formation of an adduct with sodium bisul®te. Lopez and Matson [72] used a hydrophilic hollow ®ber membrane contactor to carry out this enzymatic transformation. Their module contained ®bers made from a polyacrylonitrile-based copolymer with good solvent resistance potted in epoxy and housed in stainless steel. The interior surface of the hollow ®bers was covered by an ultra®ltration-type skin having a nominal molecular weight cutoff of about 30K, too small for the enzyme to pass, while the outer surface was open and readily permeated by the enzyme. The water-soluble enzyme was immobilized inside of these asymmetric, aqueous-®lled pores, retained by the 30K skin on the inner surface and the aqueous/ organic interface at the outer one. A solution of the ()-trans-MMPG in toluene ¯owed through the shell, while an aqueous solution of sodium bisul®te passed through the tubes. The (2S-3R)-methoxyphenylglycidic acid resulting from the enzymatic conversion was extracted into the aqueous phase, where it was immediately stabilized by the sodium bisul®te. Lopez and Matson [72] explained that their approach easily accommodates the inevitable loss of enzyme activity. Periodically, the enzyme can be removed by back¯ushing water from the tube to the shell side, then reloaded by ultra®ltering enzyme-

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containing solution in the opposite direction. Unlike covalent immobilization schemes, this approach permits easy replacement of the enzyme without the need to replace the more expensive membrane, and it does not require development of speci®c coupling chemistries. As a result of this work, a commercial scale membrane reactor system in Japan is currently producing over 75 metric tons per year of 99% (2R-3S)trans-MMPG, which accounts for over half of US Cardizem sales. The plant uses two banks of 12 modules for a total interfacial area of 1440 m2, and achieves a productivity of over 75 kg/year m2. Hoq et al. [50] also employed a membrane contactor to carry out a lipase-mediated reaction, i.e., hydrolysis of triglycerides. These researchers passed olive oil and an aqueous solution of lipase and glycerol (included to stabilize the lipase) on opposite sides of a microporous polypropylene ¯at sheet membrane with oil in the pores. The enzyme was adsorbed onto the membrane surface, and the reaction took place there. Reaction products, glycerol and fatty acids, partitioned into the aqueous and oil phases, respectively, so that the reaction and separation were accomplished in a single step. The extent of hydrolysis was measured by the amount of glycerol in the exit aqueous stream. At an oil phase residence time of about 3.6 h, the extent of hydrolysis was found to increase linearly with interfacial enzyme concentration at low enzyme concentrations. However, at higher enzyme concentrations the extent of hydrolysis was unaffected by changes in enzyme concentration, even though the extent of hydrolysis was much less than 100%. The authors explained that the process was limited by reaction at low enzyme levels and diffusion at high levels. Extent of hydrolysis increased with increased oil phase residence time, approaching 100% at residence times of about 8 and 13 h for countercurrent and cocurrent ¯ow, respectively. Hoq et al. [50] suggested that a membrane bioreactor for fat hydrolysis would offer numerous advantages over conventional emulsion systems. They also suggested that a hollow ®ber con®guration might be preferred for industrial implementation because the interfacial area is higher than that of the ¯at plate geometry they used. Similar work was reported by Molinari et al. [78], who used a 50K polyamide-based ultra®ltration membrane reactor with olive oil and water ¯owing in cocurrent mode on the shell and tube sides, respec-

tively. Candida cylindracea lipase was immobilized onto the membrane by passing an aqueous solution from the outside sponge surface to the inside skin layer of the ®ber; best results were obtained when a cross-linking agent (gluteraldehyde) was used. As with the work described above [50], diffusion through the membrane was the major resistance to mass transfer. Molinari et al. [78] compared the results from this two-¯uid reactor to those obtained with a 10K membrane reactor operated as a cross¯ow ®lter with olive oil/water emulsion fed to the tube side. The speci®c enzyme activity was over three times higher for the two-¯uid reactor (0.529 vs. 0.170 mmol/mg h). As with Hoq et al. [50], the two-¯uid arrangement allowed simultaneous reaction and product removal, relieving product inhibition and improving productivity. Furthermore, the membrane reactor is less energy intensive than traditional fat-splitting processes, which typically use an inorganic catalyst operated at temperatures of 150±2608C and pressures of 1.2± 5.0 MPa. Several industrially important organic acids produced by fermentation have been subjects of membrane contactor studies, including acetic [29,34,61,89, 91±93], lactic [20], succinic [91±93,108] and citric acid [9]. For example, Coelhoso et al. demonstrated the extraction of lactate from a 60% aqueous solution of sodium lactate using a quaternary ammonium salt carrier, Aliquat 336 (Fluka, Germany). These researchers found that the membrane resistance controlled, and resistance could be reduced by lowering the viscosity of the organic ¯uid via an increase in temperature or reduction in concentration. Simultaneous extraction and stripping of lactate (the latter with 1 M sodium chloride) was achieved by using two hollow ®ber modules in series, with the organic ¯uid circulating from one to the other and back. The use of membrane contactors in biological waste treatment is discussed in Section 6.8. 6.6. Protein extraction This topic has received increased attention in recent years with advances in genetic engineering, protein engineering and related technologies. With progress in these areas the development of economical means of protein puri®cation has become the focus of many

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investigations. Challenges include low concentration, abundance of molecules similar to the desired one, and sensitivity to heat, extremes of pH, and organic solvents. Proposed solutions include two-phase aqueous extraction [67] and reversed micelles formed from an alkane solution of oil-soluble detergent [44]. Dahuron and Cussler [29] studied protein extraction in membrane contactors using both of these approaches. Solutions of cytochrome-c, myoglobin, a-chymotrypsin, catalase, and urease in phosphate buffer were extracted using an immiscible aqueous phase, PEG. Additionally, cytochrome-c was extracted from phosphate buffer by forming reversed micelles in aerosol OT dissolved in octane. For each experiment, ¯uids were circulated from their respective reservoirs through the membrane contactor and back to the reservoirs, and the reservoir concentrations were monitored as a function of time. Equations were presented which predicted that ln(
C/
Co) would change linearly with time, and the experimental data supported that prediction. The mass transfer coef®cients were calculated from the slopes of ln(
C/
Co) vs. time plots, and values ranged from 10ÿ3 to 10ÿ6 cm/s. This is the range commonly observed for small solutes [27,133], which implies the absence of any special mechanism for protein transport. In particular:  There was no ultrafiltration effect. If there were, one would expect the lowest mass transfer coefficient for urease since it was the largest molecule studied (MWˆ250 000), yet it exhibited the largest coefficient.  There was no interfacial resistance with the inverted micelle system, an observation that can be explained by considering the mechanism of extraction using reversed micelles. The protein diffuses through the water to the fluid/fluid interface, reacts to form the reversed micelle, then the micelle diffuses through the membrane pore. The mass transfer coefficient would be reduced (vs. what is typically observed for small solutes) if the formation of the micelle were slower than diffusion through the pore, yet no such reduction was observed. Some of the challenges encountered in this work were as follows:  With the inverted micelle system the surfactant was adsorbed onto the membrane surface, which

95

lowered the interfacial tension and led to emulsion problems. Emulsions were also encountered with the two-phase aqueous system as a result of the low interfacial tension between the fluids. These problems were addressed by careful control of the static pressure difference across the membrane.  Protein denaturation (and consequent precipitation) occurred with the inverted micelle system, an issue that was addressed by taking care to remove trapped air beforehand.  In correlating the measured mass transfer coefficients to the usual dimensionless groups (i.e., Re, Gz and Sc), diffusion coefficients of proteins were difficult to obtain. Best results were obtained when diffusion coefficients in water were used. Dekker et al. [33] also studied protein extraction using a membrane contactor. Speci®cally, they employed microporous polypropylene hollow ®bers to investigate extraction of a-amylase using reversed micelles composed of iso-octane and surfactants. These investigators found that the system was stable over only a small range of differential pressure (8± 20 cm H2O), which could present dif®culties upon scale-up. 6.7. Pharmaceutical applications Prasad and Sirkar [95] characterized the simultaneous extraction of MT and CNT from an actual process stream provided by Merck (see Section 4) using a hydrophobic membrane and either toluene or benzene as the extraction solvent. Solute recoveries varied with ¯uid ¯ow rates as expected, and values >99% were reported. As mentioned earlier, surprisingly small HTU values (i.e., 3±15 cm) were obtained, and those values decreased with decreasing ®ber diameter. The tube side mass transfer coef®cient was predicted with reasonable accuracy using the Graetz equation [46,47]. No deterioration in membrane contactor performance was observed after 64 days of continuous operation. Contacting schemes using series±parallel arrangements of multiple modules were also discussed. Prasad and Sirkar [94] studied the puri®cation of mevinolinic acid also using an actual commercial stream provided by Merck. Mevinolinic acid, also known as MK-819, is the precursor to the hypocholesteremic agent mevinolin [4]. MK-819 is produced

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by fermentation of Aspergillus terreus; the product is easily extracted into isopropyl acetate, which in turn is extracted at high pH into water, and then backextracted into isopropyl acetate at low pH. Prasad and Sirkar demonstrated the ef®cacy of hydrophobic ¯at sheet and hollow ®ber membrane contactors for the latter two extractions; their work included investigation of pressure drop, breakthrough pressure, mass transfer behavior, gel formation, and cost. Several aspects of the work are noteworthy:  Pressure drop through the hollow fiber module was negligible under the conditions studied, i.e., fiber diameter 240 mm, fiber length 15 cm, and tube side velocities as high as 1 cm/s.  Acidification prior to the low pH extraction would have led to the formation of a troublesome gel phase. This was avoided by controlled addition of acid while the aqueous phase was within the hollow fibers in contact with the extractant.  The high and low pH extractions represented low and high mi systems, respectively. This means that the organic phase resistance was higher than that of the aqueous phase with the former, so a hydrophilic membrane (wetted by the aqueous phase) would have been preferred. Unfortunately, this was not feasible due to the formation of an unstable precipitate which would have blocked the aqueous phase-filled pores.  Contrary to Eq. (6a), when 3% isopropanol was added to the aqueous phase prior to the high pH extraction, the breakthrough pressure increased (from 3±4 to 7 psi at pH 12) even though the interfacial tension dropped. The authors explained these and other results in terms of the membrane pore size, the change in charge of MK-819 with pH, and the surfactant properties of MK-819. 6.8. Wastewater treatment Recovery of troublesome pollutants from aqueous solution by liquid/liquid extraction has been investigated extensively. The list of pollutants studied includes phenol, toluene, 2-chlorophenol, benzene, chloroform, 1,1,2-trichloroethane, trichloroethylene, carbon tetrachloride, tetrachloroethylene, nitrobenzene, trichloromethane, tetrachloromethane and acrylonitrile, with MIBK, hexane, isopropyl acetate,

kerosene, silicone oil and sun¯ower oil as extraction solvents [6,91,93,96,152,153]. Air stripping of VOCs has also been demonstrated. Semmens et al. [115] used microporous polypropylene hollow ®bers to strip chloroform, tetrachloroethylene, carbon tetrachloride, 1,1,2-trichloroethane and trichloroethylene from aqueous solution, with the water and organic-free air ¯owing countercurrently on the tube and the shell sides, respectively. Not surprisingly, the liquid ®lm resistance controlled, and in general the experimentally determined mass transfer coef®cient agreed reasonably well with the value predicted by Eq. (16). Semmens and Gantzer [114] discussed both the theoretical and practical aspects of VOC removal by gas stripping. Several studies have been done on the use of membrane contactors for bubble free aeration in wastewater treatment. Such an approach offers several advantages over conventional methods including absence of foaming and higher aeration rates [25,26]; additional advantages are discussed below. CoÃte et al. [26] studied the problem using a bundle of silicone rubber hollow ®bers immersed in a chamber containing water that had been treated with sodium sul®te to remove the oxygen. The temperature-controlled water was circulated across the bundle while gas (oxygen or air) ¯owed through the ®bers, and the oxygen concentration in the water was measured over time. Data were collected for a variety of conditions, and the corresponding mass transfer coef®cients were calculated; results were predicted by a model that considered two resistances in series, i.e., the membrane and the liquid ®lm (see Eq. (33)). The authors compared operation in ¯ow-through vs. dead end mode on the gas side, and reported a 45% lower oxygen transfer rate with the latter. They attributed this to the fact that in dead end mode the oxygen had to diffuse through water vapor, while with ¯ow-through operation the water vapor was swept away. Ahmed and Semmens [1] also looked at bubbleless aeration with ®bers operated in the dead end mode. These workers sealed one set of ends of a microporous polypropylene ®ber bundle, then placed the bundle inside a PVC pipe. The ®bers were connected to an oxygen supply, while water was circulated through the pipe from a reservoir, causing the ®bers to ¯uidize. The oxygen concentration in the reservoir was mea-

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sured vs. time, and from those data the overall mass transfer coef®cient was calculated. The authors concluded that liquid ®lm mass transfer coef®cients were as high as any previously reported for parallel ¯ow modules, primarily because of the high Reynolds numbers achievable with the module design employed. The mass transfer coef®cient increased with increasing pressure, but this turned out to be an artifact caused by back diffusion of nitrogen from the water into the ®bers [2]. The results indicate that one can realize the advantages of dead end operation without sacri®cing mass transfer ef®ciency. These advantages include no loss of VOCs to the gas stream and 100% oxygen transfer ef®ciency at low power input. Low power input is important in aquaculture because trout, salmon and other ®sh that require high levels of dissolved oxygen are troubled by the high power inputs associated with conventional aeration equipment [111]. In related work, Johnson et al. [53] studied bubblefree aeration using a sealed hollow ®ber bundle in a jet stream. That is, the bundle was unencumbered by a shell or similar enclosure, which allowed the bundle to become ¯uidized by the turbulent ¯ow of the jet. Johnson et al. pointed out two important advantages of this con®guration: reduced head loss, and ability to process solids-laden streams without clogging. The authors showed that these advantages are available with no reduction in mass transfer rate compared to conventional designs (see Eq. (35)). Semmens et al. [111] described ®eld studies used to test a microporous polypropylene membrane contactor operated in dead end mode. Short ®bers (less than 0.7 m) performed better than long (3 m) ones because the former moved more freely and were less likely to trap ¯occulent solids. Uncoated ®bers performed well in the aeration of secondary ef¯uent, but they fouled easily upon treatment of lake water. However, a gas permeable coating allowed the ®bers to run for over 300 h with no decrease in oxygen transfer ef®ciency. Related reports were given by Semmens and Gantzer [113], Pankhania et al. [83], and Semmens and Gantzer [114]. McGrath and Ergas [75] demonstrated the removal of toluene from air and subsequent biological destruction using a microporous polypropylene hollow ®ber module. An aqueous solution containing a mineral media inoculated with a culture of Pseudomonas

97

putida was circulated through the shell side, and a syringe pump was used to continuously add toluene to the air that ¯owed on the tube side of the ®bers. Over 98% toluene removal was achieved at an inlet concentration of 100 ppm and an overall residence time of less than 1 min. Aziz et al. [6] studied the biodegradation of trichloroethylene in a microporous polypropylene hollow ®ber contactor. They passed contaminated water (containing up to 709 mg/l trichloroethylene) through the tube side of the ®bers while circulating Methylosinus trichosporium OB3b, a methanotrophic bacterium, through the shell side. Some 78.3±99.9% of the trichloroethylene was removed from the ®bers at hydraulic residence times of 3±15 min; biodegradation rate constants ranged from 0.16 to 0.9 l (mg TSS)ÿ1 dayÿ1, where TSS is total suspended solid. The authors concluded that the hollow ®ber membrane bioreactor is a promising technology for degradation of chlorinated solvents. 6.9. Metal ion extraction There is a growing need for cost-effective technologies for recovery of metals from industrial process streams, not only because the metals are valuable, but also to meet increasingly stringent regulatory requirements. Several examples of metal recovery using membrane contactors have been reported. Yang et al. [148] described the use of membrane contactors for simultaneous extraction of cations (e.g., Cu, Zn, Ni) and anions (e.g., Cr(VI), Hg, Cd) from metalcontaminated wastewaters. These investigators used a hollow ®ber contactor with two sets of ®bers containing a liquid ion exchanger and a tertiary amine solvent, respectively; details were described by Carroll et al. [18]. The wastewater stream ¯owed on the shell side of the module, and each extractant ¯owed through one of the ®ber sets. Taking a divalent cation as an example, the extraction proceeded according to the following reaction: A2‡ …aq† ‡ 2RH …org† $ R2 A …org† ‡ 2H‡ …aq† (47) In this reaction the proton release caused an undesirable shift in the partition coef®cient towards the wastewater. However, a synergy was created by simultaneously extracting of an anion, a process that

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consumed protons: HXÿ …aq† ‡ H‡ …aq† ‡ 2R3 N …org† $ …R3 NH†2 X …org†

(48)

A similar approach was used for enhanced ef®ciency in the simultaneous recovery of two cations, e.g., Cu2‡ and Zn2‡. These cations are typically recovered using extractants like LIX 84 (Henkel, Tuscon, AZ) and bis(2-ethylhexyl)-phosphoric acid (D2EHPA; Sigma, St. Louis, MO), but selectivity is usually poor when the cations are extracted simultaneously. A signi®cant improvement in selectivity was obtained from the synergy offered by the geometry described here, with the wastewater on the shell side, LIX 84 ¯owing through one set of ®bers and D2EHPA ¯owing through the other set. Yun et al. [151] also studied extraction of copper and chromium(VI) from water using a hollow ®ber membrane contactor, with water and extracting solvent on the tube and shell side, respectively. Copper was extracted using LIX 84 in accordance with Eq. (47), while chromium(VI), present as HCrOÿ 4, was extracted using tri-n-octylamine as described by Eq. (48). Metal concentrations were reduced to 1 mg/l from 500 and 100 mg/l for copper and chromium(VI), respectively. The authors presented a mathematical model that predicted the experimental data well. They concluded that the extraction of copper was controlled by interfacial kinetics; the equilibrium constant for the interfacial reaction (Eq. (47)) is 1.7, and the forward reaction rate constant is 9.010ÿ6 cm/s. Yoshizuka et al. [150] characterized the extraction of palladium(II) from aqueous chloride media into 3,3-diethylthietane (DETE) using a hollow ®ber module. They pointed out that reactions accompanying extraction of metals (e.g., Eqs. (47) and (48)) take place either at the aqueous/organic interface or on the aqueous side close to the interface. The latter occurs when the extractant has signi®cant solubility in water, such as with chelating agents like diphenylthiocarbazone and benzoylacetone, while the reaction is con®ned to the interface with extractants such as tributylphosphate due to their long hydrophobic tails. Yoshizuka et al. showed that DETE, which is highly soluble in water, does not show interfacial activity. Equilibrium and kinetic studies allowed extraction

rates to be explained by both a diffusion model and homogeneous reaction kinetics in the aqueous phase taking into account the velocity distributions on both the shell and tube sides. Kathios et al. [57] investigated the feasibility of microporous polypropylene hollow ®ber membrane contactors for processing radioactive wastes. They chose neodymium as a model solute; this element is not radioactive but behaves similarly to trivalent americium, a long-lived actinide. The organic extracts DHDECMP (dihexyl-N,N-diethylcarbamoylmethylphosphonate) and CMPO (n-octyl(phenyl)-N,N-diisobutylcarbamoylmethylphosphine oxide) were chosen as organic extractants because of their selectivity for actinides over ®ssion products and their af®nity for trivalent actinides at moderate acid concentrations; diisopropylbenzene (DIPB) was used as a diluent. The extraction proceeded in the presence of 2 M nitric acid via the formation of a neodymium complex: Nd3‡ ‡ 3NOÿ 3 ‡ 3CMPO $ Nd…NO3 †3
3CMPO (49) (Note that Eq. (49) is similar to Eq. (27).) The reverse reaction describes the subsequent recovery of neodymium from the organic phase by back-extraction into dilute (0.01 M) nitric acid. Aqueous and organic phases ¯owed on the tube and shell side, respectively, in countercurrent mode. High organic and low aqueous ¯ow rates gave best results for the forward extraction, while low organic and high aqueous ¯ow rates were best for the back extraction. In each case the low ¯ow rate of the soluterich stream provided a long residence time for mass transfer, while the high ¯ow rate of the extracting ¯uid maintained the maximum driving force by keeping the solute concentration low. In accordance with Eq. (3), a higher extent of neodymium removal was obtained with a system having a higher value of the distribution coef®cient mi. That is, the forward reaction was increasingly ef®cient for 0.25 M DHDECMP, 0.81 M DHDECMP and 0.25 M CMPO; respective mi values were 0.25, 3.6 and 8.8. Neodymium recovery was better in the back-extraction with a low mi system, where the higher mass transfer resistance (see Eq. (3)) is more than compensated by the greater equilibrium concentration driving force. The authors suggested that a hydrophilic mem-

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brane might be more appropriate in this application. In that case the pores would be wetted by the aqueous rather than the organic phase; this would lead to a lower mass transfer resistance because the neodymium would travel in its compact ionic form rather than as a bulky complex. Additional examples of metal ion extraction were provided by Prasad and Sirkar [96]. 6.10. Semiconductors Membrane contactors are used in the production of ultrapure water for semiconductor manufacturing, where remarkably low levels of contaminating gases are required. For example, oxygen concentration must be reduced to the ppb range to avoid uncontrolled native silica oxide growth in wafer-immersion systems. Unlike conventional deoxygenation approaches like such as nitrogen bubbling and vacuum degasi®cation, membrane contactors provide a uniform water dispersion and they are insensitive to changes in ¯ow rate. Nitrogen bubbling has another disadvantage, i.e., even though it removes oxygen, it saturates the water with another troublesome gas, nitrogen. Membrane contactors are also used to add CO2 to water to increase the aggressiveness of a rinse step [31]. As mentioned in the above discussion of commercially available modules (Section 2), membrane contactors are used commercially for ozonation of water in silicon wafer cleaning applications. Speci®c uses include removal of organic and metal contamination, control of native oxide growth, and photoresist stripping. Unlike conventional approaches, this method of ozonation is bubble-free, an important feature in the semiconductor industry because bubbles attract particles that later form troublesome deposits on wafer surfaces. Membrane contactors also allow manufacturers to enjoy the advantages of ozone in lieu of conventional cleaning chemicals like sulfuric acid and peroxides. That is, ozone is milder, more environmentally friendly, reduces water usage, and requires less ¯oor space to handle. Membrane contactors have become the technology of choice for ozonation in Japan [17,142±144]. Wikol et al. [143] investigated the ozonation of tap water using a microporous PTFE hollow ®ber module (DISSO3LVE, W.L. Gore & Associates, Elkton, MD; see Section 2) with water and gas on the tube and shell

99

sides, respectively. Inlet gas ozone concentrations as high as 235 g/m3 were obtained by feeding two standard liters per minute of oxygen and 10 standard cm3/min of nitrogen to an ozone generator. The authors found that the dissolved ozone concentration in the liquid leaving the membrane contactor increased linearly with inlet gas ozone concentration and with pressure in accordance with Henry's law. They also found that ozone concentration decreased with increasing ¯ow rate of water; this indicates that increased residence time for mass transfer was more important than decreased tube side boundary layer resistance or increased concentration driving force (see Eq. (9)). When operated at a total pressure of 1.0 kg/cm2, temperature of 258C and feed gas ozone concentration of 235 g/m3, target dissolved ozone concentration of 2±10 ppm (suitable for most wafer cleaning applications) was obtained at water ¯ow rates as high as 15 l/min. Similar results were reported by Gottschalk et al. [45]. In related work, Cornelissen et al. [23] studied ozonation of degasi®ed, deionized water using W.L. Gore's DISSO3LVE PTFE membrane contactor, again with liquid and gas on the tube and shell sides, respectively. Lower oxygen ¯ow rates led to higher ozone concentrations in the gas exiting the ozone generator; the corresponding increase in mass transfer driving force led to higher dissolved ozone concentrations in the water leaving the membrane module. The extent of ozonation also increased with decreasing water pH, where base-catalyzed decomposition is less important, and with decreasing temperature, where the equilibrium ozone concentration is higher. Dissolved ozone levels increased with decreasing water ¯ow rate in agreement with the ®ndings of Wikol et al. [143] discussed above; again target ozone concentrations were obtained at a water ¯ow rate of 15 l/min, suitable for most commercial applications. 6.11. Osmotic distillation Osmotic distillation (OD) is a process where one or more volatile components of a liquid feed are transferred through a non-wetted microporous membrane into another liquid capable of absorbing those components. The driving force is the difference in vapor pressures of the volatile component(s) over the respective liquid solutions. The process proceeds at ambient

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temperature, which is particularly attractive for heat sensitive products such as those found in the food and pharmaceutical industries. For many applications of practical interest, the feed is a dilute aqueous solution of non-volatile solutes such as sugars, polysaccharides, amino acids, proteins, or carboxylic acid salts, and the objective is to concentrate the solution by removing the solvent water. A concentrated aqueous salt solution (brine) is employed on the opposite side of the membrane to provide the vapor pressure driving force; the diluted salt solution leaving the membrane contactor is concentrated then reused. A number of salts are suitable, but potassium salts of ortho- and pyrophosphoric acid offer several advantages, including low equivalent weight, high water solubility, steep positive temperature coef®cients of solubility, and safe use in foods and pharmaceuticals [49,125,126]. Osmotic distillation can also be used to concentrate aqueous solutions of volatiles by selective removal of water, a feature of interest to the ¯avor and fragrance industry. Flavor and fragrance compounds can be concentrated in this manner because (1) at the low operating temperatures typically used, the vapor pressure of these compounds is depressed relative to that of water, (2) many ¯avor/fragrance volatiles are hydrophobic, rendering them poorly soluble in concentrated saline solutions, and (3) the molecular weights of these compounds are substantially higher than that of water, so that their diffusive permeability through the membrane is much lower [49]. Fruit and vegetable juice concentrations have been successfully demonstrated at pilot facilities in Melbourne and Mildura, Australia. The Melbourne plant contains two 19.2 m2 Liqui-Cel modules (see Fig. 1), along with ultra®ltration and reverse osmosis equipment (for feed pretreatment) and a brine evaporator. The plant can concentrate 50 l/h of juice to 65±70% solid by weight. The Mildura plant is operated by Vineland Concentrates (a subsidiary of the Wingara Wine Group) and CELGARD LLC. This plant contains 22 4 in.28 in. Liqui-Cel modules (total interfacial area: 425 m2) operated at 30±358C and two atmospheres. Color, ¯avor and aroma retention are good, but module cleaning is an issue [49,86,125,126]. Varietal grape juice concentrates used for production of high-quality wines are of particular interest. These concentrates are stable upon long periods of

storage; as a result, they can be used as needed to adjust the sugar content of fresh grapes so that variations in the alcohol content of a particular wine are minimized. Furthermore, such concentrates can be shipped over long distances to allow production of high-priced wines in regions where the required grapes are not available or too expensive [49]. Tomato juice concentration has also been studied in recent work (e.g., [37,87]). The growing emphasis on health and ®tness in recent years has led to a demand for beer and wine with reduced alcohol content. Such products can be prepared by osmotic distillation if water is used as the stripping liquid instead of saline solution. This leads to selective removal of ethanol without signi®cant depletion of water or the components that give beer and wine their characteristic taste, aroma and mouthfeel. Similarly, OD can be used to reduce the alcohol content of ``hot'' wines, i.e., wines whose alcohol content is too high because they are produced from grapes with high sugar. J. Lohr Winery (San Jose, CA) is currently using four 4 in.28 in. Liqui-Cel modules to produce Ariel, a wine that contains less than 0.5% alcohol [49,76,86]. In the pharmaceutical area, OD can be used to concentrate miscible solvent extracts of intracellular products from fermentation broths. Similarly, solvent can be selectively removed from chromatography column eluants to concentrate antibiotics, vaccines, hormones and other heat-sensitive biologicals [49]. Osmotic distillation has been recently reviewed by Kunz et al. [68] and Hogan et al. [49]. 7. Concluding remarks Membrane contactors embody an exciting new approach to ¯uid/¯uid contacting that offers numerous advantages, including substantially higher ef®ciencies (i.e., lower HTU values) than those achievable in traditional dispersed phase devices. Membrane contactors have been demonstrated in a diverse range of gas/liquid and liquid/liquid applications; even though several of these have been commercialized, the number of commercial applications has not yet reached the full potential of the technology, particularly with liquid/liquid contacting. To facilitate further commercialization, challenges that need to be addressed

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include (1) better understanding of shell side ¯ow leading to more reliable scale-up procedures, (2) higher comfort level with the technology among potential users, and (3) improved materials of construction that offer a wide range of solvent compatibility. Undoubtedly, the advantages of membrane contactors will render the technology pervasive as progress is made in these areas. 8. Nomenclature a A C Cp d de D f G Gz h H HTU k k0 hki K KE kR L LTU m MG MT NTU Nu Pr Q r R0 Re S

interfacial area (cm2/cm3) membrane area in module (cm2) concentration (g mol/l); constant in Eq. (25) heat capacity at constant pressure (J/s K) diameter (cm) (mm in Figs. 11±14) equivalent diameter (cm) diffusivity (cm2/s) function of geometry in Eq. (15) gas flow rate (cm3/s) Graetz number, d2v/DL heat transfer coefficient (J/s cm2 K) Henry's law constant height of a transfer unit (cm) individual mass transfer coefficient (cm/s); thermal conductivity (J/s cm K) mass transfer coefficient with instantaneous reaction (cm/s) average mass transfer coefficient (cm/s) overall mass transfer coefficient (cm/s) equilibrium constant (g mol/l) reaction rate constant (sÿ1) length (cm) length of a transfer unit (cm) partition coefficient: ratio of equilibrium concentration in organic phase to that in aqueous phase module geometry number (rr/zz)(L/R)2 mass transfer number, kA number of transfer units Nusselt number, hd/k Prandtl number, Cp/k flow rate (cm3/s) radius (cm) average radius (cm) Reynolds number, dv/ cross-sectional area (cm2)

Sc Sh hShi v V

101

Schmidt number, /D Sherwood number, kd/D average Sherwood number, hkid/D velocity (cm/s) average volume occupied by one fiber (cm3)

Subscript A b B E e f G i i in L lm M Nb o out r s t w z 1 0

component A bulk component B effective; equilibrium (Eq. (46)) equivalent feed gas component i inside; at the interface module inlet liquid log mean membrane number of baffles organic phase; outside; initial condition module outlet radial shell side tube side aqueous phase; wall axial end of module where organic phase exits end of module where organic phase enters

Greek letters   0 0

    

exponent for Reynolds number in Eq. (15); extent of shell side bypassing exponent for Schmidt number in Eq. (15); constant in Eq. (19) membrane porosity void fraction of fibers standard deviation divided by the mean (Eqs. (17) and (29)) interfacial tension (dynes/cm) fiber packing fraction Darcy's permeability tensor viscosity (g/cm s) contact angle density (g/cm3)

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surface tension (dynes/cm) space time (s); tortuosity

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