Novel Membrane and Device for Direct Contact Membrane Distillation

In the direct contact membrane distillation (DCMD) process for desalination, the ...... David M. Warsinger , John H. Lienhard V , Mikel C. Duke , Take...
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Ind. Eng. Chem. Res. 2004, 43, 5300-5309

Novel Membrane and Device for Direct Contact Membrane Distillation-Based Desalination Process Baoan Li and Kamalesh K. Sirkar* Otto H. York Department of Chemical Engineering, New Jersey Institute of Technology, Newark, New Jersey 07102

In the direct contact membrane distillation (DCMD) process for desalination, the water vapor flux is strongly affected by the hot brine heat transfer coefficient, conductive heat loss, and long-term flux decline due to membrane pore wetting/fouling, etc. The DCMD process has been explored here using porous hydrophobic polypropylene hollow fibers having three different dimensions and two different wall thicknesses. The outside surfaces of the fibers have been coated with a variety of microporous plasmapolymerized silicone-fluoropolymer coating. A large number of rectangular modules having the hot brine in cross flow over the outside of the fibers and cold distillate flowing in the tube side have been investigated for their DCMD performances with hot brine (1% NaCl) over a brine temperature range of 60-90 °C. The module MXFR 3 containing fibers with larger internal diameter (i.d.) and wall thickness, and having the best performance, was tested in a continuous DCMD run for 120 h with 85 °C brine flowing at a Reynolds number of 70. The remarkably high water vapor flux (41-79 kg/m2‚h) obtained in such modules and the complete absence of pore wetting over 400 h of experiments without any module cleaning demonstrate the excellent DCMD potential of such hollow fiber membranes and modules. 1. Introduction In direct contact membrane distillation (DCMD) of an aqueous solution containing a nonvolatile solute, e.g., brine, a porous hydrophobic membrane is imposed between the aqueous solution at a higher temperature and a colder distillate stream on the other side of the membrane (Figure 1a). Water vapor evaporated from the hot brine (solution) diffuses through the porous membrane and condenses into the cold distillate on the other side. The higher the temperature of the hot solution (brine) at the membrane-solution interface, the higher the water vapor pressure and therefore the higher the water vapor flux through the membrane. The heat for evaporation of water is supplied from the sensible heat of the hot feed solution. The earliest reports of membrane distillation are available in Findley1 and Gore.2 Notable earlier additional research was carried out by Schofield et al.3 and Schneider et al.4 A brief review is available in Sirkar;5 a more extensive review of membrane distillation has been provided by Lawson and Lloyd.6 These earlier studies and reviews point out the important role of heat transfer coefficients in achieving high water vapor flux in membrane distillation. A high feed side heat transfer coefficient is needed to reduce the temperature drop in the hot feed from the bulk to the membrane surface, i.e., to reduce the temperature polarization (Figure 1b). The closer Tfm is to Tf, the higher is the water vapor pressure at the hot brine-gas interface. Successful membrane performance in DCMD therefore requires a high feed side heat transfer coefficient;3,5-12 the requirement is generally less stringent on the distillate side because the temper* To whom correspondence should be addressed. Tel.: 1-973596-8447. Fax: 1-973-642-4854. E-mail: [email protected] (K. K. Sirkar); [email protected] (Baoan Li).

ature coefficient of water vapor pressure is quite low at the lower distillate temperatures. As water vapor is transferred through the porous membrane, heat is also transferred from the hot brine to the cold distillate via conduction through the solid polymer as well as through the vapor space in the pores. This heat flow is a loss insofar as evaporating water and recovering it is concerned. Therefore, conductive heat transfer needs to be reduced to improve DCMD performance. Because thermal conduction through the vapor space is usually much less than that through the solid polymeric wall, a more porous membrane and a higher membrane wall thickness will reduce the extent of heat loss by heat conduction. Membranes having a higher porosity and used in DCMD often have a larger pore size (0.1-0.6 µm). Hydrophobic membranes of these larger pore sizes used in DCMD are not to be wetted by aqueous nonwetting solutions such as saline water under the usual operating conditions unless the solution pressure exceeds the breakthrough pressure (∆Pbr).4,5 Yet the observed water vapor flux at pressures much below ∆Pbr under these conditions suffers flux decay with time. This is hypothesized to be primarily due to pore wetting as well as due to pore fouling, scaling, etc.4,6,13-16 Membrane design which prevents/controls these phenomena is of significant interest. Hollow fiber-based membrane devices for DCMD are likely to be of considerable utility because they are simple, potentially scalable, and can often pack a large membrane surface area per unit device volume without any need for a supporting structure. Usually in such a device, the distillate will flow in the tube side. Excessive heating of the distillate in DCMD will be counterproductive since it can reduce the temperature driving force for water evaporation across the membrane pore. Hollow-fiber module design vis-a`-vis the distillate flow rate

10.1021/ie030871s CCC: $27.50 © 2004 American Chemical Society Published on Web 07/27/2004

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croporous silicone-fluoropolymer layer on the outside surface (Figure 1c). To reduce the water vapor transport resistance, the hydrophobic PP hollow fiber substrate pore size should be as large as possible with a high porosity. However, such pores become more susceptible to breakthrough, wetting, and fouling. It would be ideal to have a thin highly H2O vapor permeable nonporous coating over such pores. If one were to attempt to close the large substrate pores completely, the overall coating thickness would become substantial especially near the pore wall. As a result, the water vapor permeation rate will be decreased by orders of magnitude. To avoid such a result, our coating is microporous. A microporous coating of a silicone-fluropolymer will effectively increase the resistance to the fouling of the PP substrate pores due to a considerably reduced critical surface tension as well as the requirement of a much higher breakthrough pressure for aqueous systems. Second, we have employed hollow fibers of larger wall thickness and considerable porosity. Larger wall thickness and higher porosity reduce conductive heat loss while the water vapor diffusional resistance through the pores is also reduced. Third, we have increased the internal diameter of the hollow fiber to ensure a high flow rate of the cold distillate which will undergo only a low flow pressure drop. This will simultaneously ensure a lower distillate temperature rise. Fourth, and most importantly, we have implemented cross flow (Wickramasinghe et al.17 and Yang and Cussler18) of the hot brine in the module over the outside surface of the hollow fibers to drastically reduce the temperature polarization by enhancing the brine heat transfer coefficient at low brine flow rates and correspondingly, lower brine Reynolds numbers (Figure 1d). Incorporation of all such aspects in the module has led to considerably enhanced water vapor flux and high module productivity. We report these results here; further, we illustrate the results for a 120-hour-long continuous run in DCMD using the most productive membrane-based hollow fiber module. The rectangular cross-flow-based hollow fiber membrane modules employed here have surface areas varying between 113 and 257 cm2. The DCMD performances of scaled-up modules of much higher surface area (2864 cm2) will be reported later. 2. Experimental Section

Figure 1. (a) Conventional direct contact membrane distillation. (b) Temperature and partial pressure profiles in direct contact membrane distillation. (c) Water vapor flowing through microporous hydrophobic coating on the surface of porous hollow fiber membrane. (d) Cross flow over hollow fiber outside surface.

is likely to be quite important in developing practical devices/designs for a successful DCMD process. To address these four issues in DCMD, we describe here a novel hollow fiber membrane and a cross-flow module design and illustrate their performances in DCMD of hot saline water at temperatures varying between 60 and 90 °C. First, to reduce the potential for flux decay, hydrophobic porous polypropylene (PP) hollow fibers have been coated with an ultrathin mi-

2.1. Membranes and Membrane Modules. Seven cross-flow rectangular hollow fiber membrane modules (Figure 2) and one parallel flow hollow fiber membrane module have been studied. Polypropylene (PP) porous hollow fiber membranes in these modules were from Accurel MEMBRANA (Wuppertal, Germany); they are identified as PP 50/200, PP 50/280, and PP 150/330. All fibers were coated with different proprietary recipes by plasma polymerization at Applied Membrane Technology (AMT) Inc. (Minnetonka, MN) except the fibers in module MXFR 15. Physical properties of the hollow fibers and characteristics of the membrane modules are listed in Table 1. The hollow fibers were arranged in a staggered fashion in the rectangular modules. The liquid entrance on the shell side to the rectangular module had a well designed diverging section; the exit similarly had a converging section (Figure 3a). These potentially allowed the liquid (hot brine) to flow uniformly in cross flow outside of and perpendicular to the fibers in the modules. The diverging section and the

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Figure 2. Rectangular cross-flow test module without face plates. Table 1. Details of the Hollow Fibers and the Membrane Modules Used particulars

MXFR 6

support membrane type fiber i.d., µm wall thickness, µm maximum pore size, µm membrane porosity coatinga no. of fibers (no. of layers × fibers/layer) effective membrane surface area, cm2 c packing fraction shell side flow mode module frame (internal dimensions)

PP 50/200 200 52.5 0.1 0.50

MXFR 9

MXFR 10

MXFR 13

PP 50/280 280 50 0.1 0.50 silicone fluoropolymerb 624 456 (13 × 48) (12 × 38) 251 256.6 0.10

MXFR 15

uncoated

MXFR 1

MXFR 3

module 4

PP 150/330 PP 50/200 330 200 150 52.5 >0.20 0.1 0.65 0.50 silicone fluoropolymerb 268 180 300 (2 × 20+12 × 19) (10 × 18) 178 119 491

0.13 0.18 rectangular cross-flow rectangular: length 6.4 cm, width 2.5 cm, height 1.8 cm

0.12

0.30 parallel tube: i.d. 1 cm; length 17 cm

a All coatings were applied on the outside diameter of the support fibers by Applied Membrane Technology, Inc., Minnetonka, MN using their proprietary plasmapolymerization technology. b Each of the modules of MXFR 6, Module 4, MXFR 1, MXFR 3, MXFR 9, MXFR 10, and MXFR 13 represent different coatings, respectively. However, the coating composition and preparation are closer between MXFR 1 and 3. c Based on fiber internal diameter.

Figure 3. (a) Face box fabricated for rectangular cross-flow module; (b) face plate fabricated for rectangular cross-flow module; (c) rectangular cross-flow test module with face boxes, face plates, and assembly.

converging section were two boxes with curved shapes. Two face plates, with a wide size distribution of open holes, were made from two flat plastic sheets (Figure

3b). The holes near the center were smaller and those further out were larger. The design mentioned above attempted to ensure a uniform flow of the feed solution through the shell side of the fibers. The material used for the face boxes and face plates was clear cast acrylic plastic having a reasonable thickness and heat transfer resistance. Two face boxes and face plates were assembled with a rectangular membrane module channel to constitute the complete device (Figure 3c). Neoprene gaskets were used between the face box, the face plate, and the module channel on each side to seal the parts together. Hot brine was allowed to enter one face box, leave the box through the face plate holes, which distributed the liquid flow evenly, and then enter the flow channel. On the other side, the liquid left the channel through the face plate holes and collected in the face box and then flowed beyond the box and thus the module. In this design, there was no free space between the faces of the two boxes and the fiber layer. Therefore, the liquid crossed the fiber layer perpendicularly to ensure good heat and mass transfer.

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Figure 4. Experimental setup for direct contact membrane distillation.

2.2. Experimental Details. DCMD. A schematic of the experimental apparatus is shown in Figure 4. The system piping and storage tanks were thoroughly insulated to minimize heat loss to the environment. The liquids employed in the experiments were pure deionized (DI) water on the distillate side and 1% (wt) solution of NaCl as the brine feed. In the experimental setup (Figure 4), the feed solution was introduced to the shell side from a reservoir by a digital Masterflex peristaltic pump at a constant flow rate. The connecting tubing was immersed in a thermostated water bath before the feed entered the module. A temperature controller maintained the bath temperature at a given value and thus ensured the required temperature for the hot feed. Outside the membrane module, the feed was circulated back to the feed reservoir and was rewarmed. There was continual addition of fresh water to the saline water side to maintain its total volume and salt concentration constant. Deionized water was introduced as the cooling liquid on the fiber lumen side from a reservoir by another digital Masterflex peristaltic pump at a constant flow rate. The connecting line was immersed in a ColeParmer Polystat refrigerated bath (model 12111-20) at a given low temperature before the water entered the module. The inlet and outlet temperatures of the hot feed and the cold distillate were measured by four thermocouples with 0.1 °C accuracy. The pressure drops of the feed and the distillate through the module were monitored also. The electrical conductivity or the salt concentration of the samples was measured by a conductivity meter (model 115, Orion Research, Beverly, MA). When the readings of the flow rates of the hot feed, cold distillate, and the four inlet and outlet temperatures reached constant values, it was assumed that steady-state had been reached. Subsequently, the volume increase in the cooling water reservoir was used to calculate the water permeation flux through the membrane under the given experimental conditions. Water vapor flux (NV) was calculated from the following relation:

( )

kg ) m2‚h volume of water transferred (L) × density of water (kg/L)

NV

2

membrane area (m ) × time (h)

(1)

Here, the membrane area was calculated based on the hollow fiber inside area: S ) nπdiL, where n is the

number of fibers in the module, di is the fiber i.d., and L is the fiber length. Many runs were repeated a number of times; the reproducibility in the data was (2%. Gas Permeation. A system was also established for the measurement of gas permeance of the coated porous hollow fiber membranes using a gas permeation apparatus. The N2 gas from a cylinder permeated through the membrane from the tube side to the shell side. The pressures of the upstream and downstream were measured by an Ashcroft Test Gauge (63-5631). The downstream flow rate of the gas was measured using a soap bubble flow meter. During the permeation measurements, the upstream pressure was maintained at a constant pressure, between 0.1 and 0.6 psig (0.5-3.1 cm Hg gauge). The permeation measurements were made at room temperature. The N2 permeance of the hollow fiber membranes was related to the measured steady-state permeation rate of nitrogen through the membrane by

QN2 δM

(permeance) )

P1V˙ 1T0 P0T1‚S‚∆PN2

(2)

In eq 2, T0 ) 273.15 K, P0 ) 760 mmHg, ∆PN2 corrected to STP is pressure difference across the membrane, S is the inside membrane area, P1 is the atmospheric pressure, T1 is the room temperature, V˙ 1 is the volume flow rate of gas through the membrane during measurement at room temperature, QN2 is the permeability coefficient of N2 permeation through the membrane of effective thickness δM. Calculation of Reynolds Numbers. Reynolds number is normally defined in the following way:

Re )

dvF µ

(3)

where Re is the Reynolds number, d is a characteristic dimension, v is the velocity, F is the density, and µ is the dynamic viscosity. In this paper, the Reynolds numbers of the feed or the distillate flowing through the shell or the tube side were defined as diameter-based Reynolds number (Red). In the calculation of Red based on eq 3, fiber i.d. (di) and linear velocity are used for tube-side parallel flow, and fiber o.d. (do) and interstitial velocity are used for shell-side cross-flow.

interstitial velocity ) brine flow rate/open area for flow through the shell side (4) linear velocity ) flow rate/open area for flow through the tube side (5) Here, open area for flow through the tube side ) nπ(di/2)2; open area for flow through the shell side ) frame cross sectional area (6.4 × 2.5, cm2) - fiber projected area (no. of fibers in one layer × do × L, cm2). 3. Results and Discussion We present first the results of N2 permeances of different membrane modules. Next, we report the DCMD performances of a variety of modules and critically compare their performances and select the best module. Detailed performances of the best module are then studied over a range of conditions including a

5304 Ind. Eng. Chem. Res., Vol. 43, No. 17, 2004 Table 2. Nitrogen Permeation Properties and DCMD Performances of the Hollow Fiber Membrane Modules Used particulars shell side flow mode permeance of N2,a cm3(STP)/cm2‚s‚cmHg DCMDb (below): NVc, kg/m2‚h Red (tube side) Red (shell side) pressure drop (tube side)d, kPa ηe

MXFR 6

MXFR 9

MXFR 10

MXFR 13

MXFR 15

MXFR 1

MXFR 3

module 4

0.006

0.011

0.009

cross flow 0.017

0.018

0.153

0.196

parallel flow 0.013

3.9

1.4

2.9

4.3

5.3

32.9

41.4

0.2

42 29 10.0

50 35 9.6

52 38 9.0

56 39 6.9

54 37 9.0

65 54 6.9

68 58 6.9

31 33 20.0

0.064

0.033

0.061

0.082

0.113

0.560

0.650

a

Experimental conditions: 25.5 °C, 76 cm Hg; N2 inlet, tube side; N2 outlet, shell side. b DCMD: shell side, 1% saline water at 85 °C (inlet temperature) at 200 cm/min of interstitial velocity; tube side, DI water at 15-17 °C (inlet temperature) at 760 cm/min of linear velocity. c NV water vapor flux. d Pressure drop of cold distillate on the tube side along the module length. e η: Evaporation efficiency.

continuous run over 5 days. We then consider various heat transfer related aspects of the performances for these modules in DCMD. 3.1. N2 Permeance of Different Modules. Gas permeation measurements indicated that the N2 permeance of MXFR 3 reached 0.196 cm3(STP)/cm2‚s‚cmHg which is at least 10 times higher than those of MXFR 6, MXFR 9, MXFR 10, MXFR 13, and MXFR 15 (0.0060.018 cm3(STP)/cm2‚s‚cmHg) (Table 2). Here, the values of the nitrogen permeance directly illustrate the resistance to N2 permeating through the microporous coating and the porous substrate at 22 °C ( 2 °C. The open area for nitrogen/water vapor molecule permeation not only includes the area of the support pore but also the open area of the microporous coating membrane covering the larger open pore mouth of the hollow fiber. A high substrate porosity, larger pore size, a suitable opening in the coating, and a very thin and permeable coating are important for getting very high water vapor flux in DCMD performance. Therefore, MXFR 1 and MXFR 3 are likely to have higher water vapor fluxes than the others under identical experimental conditions in DCMD. The differences between water vapor fluxes through different coatings in different modules at higher temperatures are likely to be much less than those for N2 fluxes at room temperature. 3.2. DCMD Performances of Different Modules. We provide first a comparison of the DCMD performances of different modules studied here under a particular condition, namely, 1% hot brine flowing through the shell side at an inlet temperature of 85 °C and distillate water flowing through the fiber bore side at a linear velocity of 760 cm/min. The hot brine flow velocity (calculated as the interstitial velocity between the fibers) was varied over a range of 23 to 430 cm/min. To illustrate, the Reynolds numbers for module MXFR 3 for interstitial velocity varying from 23 to 261 cm/min were 6.4 to 78.6; the corresponding values for module MXFR 6 for interstitial velocity varying from 60 to 365 cm/min were 8.5 to 53.2. The objective was to isolate those modules which have satisfactory DCMD performances. One could then explore in detail the factors which affect DCMD performances of the best membrane module as well as those of others. To that end, we compare in Figure 5a and Table 2 the DCMD performances of seven rectangular cross-flow membrane modules, namely, MXFR 1, MXFR 3, MXFR 6, MXFR 9, MXFR 10, MXFR 13, and MXFR 15 under identical conditions. Modules MXFR 9, 10, 13, and 15 employ the same PP 50/280 fibers. The fibers in module 15 do not have any plasmapolymerized coating; this module has there-

fore the highest water vapor flux among these four modules especially at higher brine-side velocities (Figure 5a). Modules MXFR 13, 10, and 9 have progressively decreasing water vapor fluxes. Generally, this decreasing performance coincides with a decrease in the N2 permeance as shown in Table 2. Because all four modules had essentially identical designs for rectangular cross-flow, one can also argue that the performance differences are most likely due to the differences in the coating resistances of the fibers in each module. However, since water vapor flux-based performances are an order of magnitude smaller than those of modules MXFR 3 and MXFR 1, it is not useful to deliberate further on these coatings and modules. It is, however, useful to compare the performance of the only parallel flow module (module 4) with that of the rectangular cross-flow module (MXFR 6) since both were prepared from the same PP 50/220 fibers. From N2 permeance values in Table 2, the coating resistance of the fibers in MXFR 6 appears to be twice that of the fibers in module 4. Yet the DCMD water vapor flux of MXFR 6 appears to be many times higher than that of module 4 (Table 2). This performance difference has now been illustrated in Figure 5b over a wide range of hot brine side velocities for a given tube-side cold distillate velocity. This huge performance difference is almost totally due to the difference in flow patterns between the two modules since they have the same substrate fibers. The two-times higher N2 permeation rate in module 4 should have a limited effect. Consider modules MXFR 10 and 13; they have identical fiber substrate and cross-flow module design. The N2 permeation rate of module MXFR 13 is about twice that of MXFR 10; the water permeation flux is, however, correspondingly larger by almost 48% (Table 2). Thus flow pattern is critical between module 4 and MXFR 6. (Note: for module 4 we have parallel flow on the shell side; therefore interstitial velocity is obtained by dividing the flow rate by the free shell-side cross-sectional area). Because brine side heat transfer coefficient has considerable influence over the water vapor flux in DCMD, rectangular cross-flow devices can easily outperform a parallel flow device, say, by orders of magnitude. At comparable flow velocities, the rectangular cross-flow can generate mass-transfer coefficients which are considerably higher.17,18 Further, the parallel flow device used here probably suffered from considerable bypassing and channeling on the shell side. Gryta and Tomaszewska19 obtained the DCMD water vapor flux of PP 150/330 hollow fiber in a parallel flow module to be 3 kg/m2‚h when feed (water) flowed through shell side at 60 °C at a velocity of 1980 cm/min and permeate was

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Figure 6. DCMD: Variation of water vapor flux with interstitial velocity of hot brine (1% NaCl) as feed flowing through the shell side (cross flow) at various temperatures (modules MXFR 3 and MXFR 1; tube side, 15-17 °C deionized water, average velocity 1660 cm/min).

Figure 5. (a) Comparison of the DCMD performances of various modules: variation of water vapor flux with velocity of hot brine (1% NaCl) as feed flowing on the shell side at inlet temperature of 85 °C (tube side, DI water, 15-17 °C, linear velocity 760 cm/ min). (b) DCMD: Variation of water vapor flux of membrane modules MXFR 6 and module 4 with velocity of hot brine (1% NaCl) as feed flowing through the shell side at inlet temperature of 85 °C (tube side, DI water, 15-17 °C, linear velocity 760 cm/min).

on tube side at 1500 cm/min. Our experiment yielded a very different result using coated PP 150/330 fibers in a rectangular cross-flow module under similar experimental conditions, namely, 22 kg/m2‚h of water vapor flux (Figure 6). 3.2.1. DCMD Performances of Module MXFR 3. Of all the modules studied and considered in Figure 5a, module MXFR 3 yielded the highest water vapor flux; for example, a value of around 45 kg/m2‚h for a brine inlet temperature of 85 °C. The performance of module

MXFR 1 was quite close. The difference in the performances of the two modules can be traced potentially to two factors: Module MXFR 1 had a significantly lower N2 permeation rate compared to that of module MXFR 3 (as shown in Table 2) indicating higher coating resistance in MXFR 1 since the substrates were identical; module MXFR 3 had a lower packing fraction. Therefore, we emphasized studying the performances of this module MXFR 3 vis-a`-vis the feed temperature and the feed velocity. We show in Figure 6 the variation of water vapor flux in the module MXFR 3 with interstitial brine velocity at four different feed brine inlet temperatures: 60, 70, 80, and 85 °C for a given tube-side velocity. The water vapor flux increases with the hot brine velocity in general, but much more so at the higher temperatures. At the higher brine-side velocity studied, brine at 85 °C yielded a flux of 69 kg/ m2‚h. One data point was obtained at 90 °C at the higher brine-side velocity of 229 cm/min through module MXFR 3; the flux was the highest at 79 kg/m2‚h. For a given brine temperature, an increase of interstitial velocity in cross flow on the shell side leads to an increase of Reynolds number which maximizes the brine-side boundary layer heat transfer coefficient. Higher heat transfer coefficient leads to higher sensible heat loss which supports the increased water vapor permeation flux. Meanwhile, the increase of feed flow rate decreases the residence time of feed in the module, and increases the feed outlet temperature which increases the vapor pressure-based driving force and the Reynolds number. Figure 7 illustrates the variation of feed outlet temperature with its interstitial velocity. It is obvious that the temperature difference between the inlet temperature and the outlet temperature becomes larger with an increase of inlet temperature at a given velocity on the shell side. That is because the high water vapor flux at a high temperature causes more heat removal from the feed solution for water evaporation.

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Figure 9. DCMD: Variation of water vapor flux and distillate outlet temperature with linear velocity of distillate (DI water) flowing through the tube side at entrance temperature of 15-23 °C (module MXFR 3; shell side, brine solution (1% NaCl) at 85 °C, average velocity 229 cm/min).

Figure 7. DCMD: Variation of feed outlet temperature with feed interstitial velocity through the shell side of hollow fiber module at various inlet temperatures (MXFR 3; feed, 1% NaCl; distillate, DI water at 15-17 °C, 1660 cm/min of linear velocity).

Figure 8. DCMD: Variation of water vapor flux with inlet temperature of hot brine (1% NaCl) as feed flowing through the shell side (cross flow) at various interstitial velocities (module MXFR 3; tube side, 15-17 °C DI water, velocity 1660 cm/min).

A replotting of the data in Figure 6 is shown as a function of brine temperature in Figure 8; the isolated data point for brine at 90 °C and the highest brine interstitial velocity of 229 cm/min yielded the highest

water vapor flux of 79 kg/m2‚h. For all experiments, the concentration of salt measured in the distillate stream was always less than 8 ppm. (It is necessary to point out here that the conductivity readings in the starting distillate stream before the experiment were around 8 ppm. So the conductivity of the distillate was lower than 8 ppm throughout the experiments). Data in the two Figures 6 and 8 illustrate how strongly the feed brine temperature and feed brine velocity in the cross-flow mode influence the water vapor flux. Correspondingly, they suggest how the water vapor flux will decrease along the hot brine flow direction in a cross-flow membrane module as the brine is cooled and the temperature is decreased. This would illustrate the diminishing returns in flux as the membrane area is increased in the brine flow direction. This may suggest that designs in which larger brine flow cross sectional area is available at high brine flow rates may be more desirable. DCMD flux data for cross-flow modules shown in Figures 5a and b, 6, and 8 were obtained at two tubeside velocities. Figure 9 illustrates how the water vapor flux in module MXFR 3 increases with an increase in the tube-side distillate velocity; this is due to an increase in the tube-side heat transfer coefficient and the corresponding decrease in Tpm. However, temperature polarization on the distillate side is not as critical; the water vapor pressure at the relatively low distillate temperature changes very slowly with temperature. Distillate velocity increase (at the cost of increasing distillate-side pressure drop) can still contribute considerably to increasing the water vapor flux. 3.2.2. An Extended DCMD Run with Module MXFR 3. All of the above data were acquired in experiments based on short-term runs (around 10 h per each run). Figure 10 shows how the water vapor flux in module MXFR 3 changed with time for a feed of 1% hot brine and cold DI water at the tube side. For a brine feed at 85 °C, the experiment was carried out over an extended period of 5 days. We wish to note that no filtration/microfiltration units were used in our system; therefore deposition of dirt, etc., on fiber surfaces and tube-sheets could not be avoided. A decreasing permeate flow rate was observed in the first 90 h of the experiment, after which a constant steady-state permeate

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Figure 10. DCMD: Variation of water vapor flux with operating time for hot brine (1% NaCl) recirculating through the shell side with a velocity of 234 cm/min at 85.5 °C (Re ) 70), and cold distillate water recirculating through tube side at a velocity of 1625 cm/min at 16 °C (Re ) 120); a data point (b) from a restarted experiment under identical experimental conditions after the extended-term experiment (module MXFR 3) (no prefiltration).

flow-rate was reached. There was a reduction of about 23% in the water permeation flux. The stable water vapor flux was 54 kg/m2‚h. Although the role of dirt, etc., depositing on the membrane surface cannot be ruled out, a partial reason for this decrease of water vapor flux possibly is a partially reversible thermal creep in the membrane and the coating material with time around the mouth of the partially covered pore at a high temperature. The phenomenon is in part reversible and initial fluxes can be restored by restarting the experiment, as shown by the data point appropriately marked. We have observed that the module MXFR 3 displayed a significantly higher DCMD flux of H2O vapor (∼60 kg/m2‚h) at the beginning of a new experiment after the 5-day-long extended-term run was over. We did not undertake any cleaning of the module before or after the extended-term run. The conductivity of the cold distillate was monitored during this extended experiment. The concentration of salt was always less than 8 mg/L, which indicates that the membrane pores were not wetted by the hot brine during this experiment. However, the pressure drop in the cold distillate water passing through the lumen side of the module was slightly increased (by 25%). This indicates the possibility of dirt buildup in the hollow fiber tube sheet. In the absence of any prefiltration, flow reversal should clean it. The membrane module MXFR 3 was used for around 400 h. During this period, it was continually used for DCMD and vacuum membrane distillation (VMD) tests (the results from the latter will be reported later). We have, however, noticed that the color of the fibers in module MXFR 3 became slightly yellowed. The yellowing is believed to have been caused by some dirt depositing on the coated section due to whatever reasons, although thermal effects cannot be ruled out. Since the bulk of the area of the open mouth of the pore was covered by a plasmapolymerized coating, the dirt deposits on the coating around the open mouth of the pores reduced the permeability of the water vapor through the coating to some extent. It is useful to deliberate on the implications of the water vapor flux levels achieved here. If we can assume that a 55-60 kg/m2‚h level water flux is achievable, then this productivity is either higher than or comparable to the values achieved in RO (25 kg/m2‚h in Ray20 and 56 kg/m2‚h in U. S. Bureau of Reclamation 21). Since the hollow fiber module design used here could be scaled up and the hollow fiber packing density is not going to

be inconsiderable, the volumetric productivity will also be high. Further, DCMD cost is intrinsically significantly lower than that of RO especially due to the lowpressure operation of DCMD, high water vapor flux, and good anti-fouling properties of the DCMD membrane when waste heat cost is taken into account.22 There now appears to be a significant basis for further scaled-up studies in DCMD. 3.2.3. Additional Observations in DCMD Experiments. During this study, it was observed that some of the fibers in MXFR 6, MXFR 9, MXFR 10, MXFR 13, and MXFR 15 came close to one another during and after the DCMD runs. Such conditions were not visible in modules MXFR 1 and 3. In the polar liquid phase of water, nonpolar fibers tend to be together to reduce the surface free energy by affinity on the surface of the hollow fibers. Two opposite factors, the packing fraction and the outer diameter of fiber, affect the extent of stickiness. Small o.d. fibers assembled with high packing fraction in module can easily come together in water. In contrast, the fibers in MXFR 3, having the larger o.d. (630 µm) and packed with a smaller packing fraction (0.12), are separated from each other. Of course, the sticking of fibers would decrease the effective surface of membrane for water permeation a little, and affect MD performance negatively. A fiber mat as currently used in Liqui-Cel modules (Celgard, Charlotte, NC) should eliminate this problem altogether. 3.3. Heat Transfer and Related Aspects in DCMD Performances of Different Modules. One of the primary concerns in DCMD, with regard to effective energy consumption, is reducing the heat lost in the process. Heat loss can be divided into two parts. Heat can be lost through the equipment surface area as the liquids flow from cooling/heating reservoirs to the membrane module. This heat loss can be minimized or eliminated by optimizing the design and insulating the equipment well. The second loss is from heat conducted through the membrane in the module from the feed to the cold distillate. This conductive flux consists of the sum of the conductive heat flux through the nonporous part of the polymeric membrane and the conductive heat flux through the gas/vapor in the pores of the membrane, which is a function of the membrane design and the thermal conductivity of the membrane material. Consequently, the loss of this conductive flux directly results in a decrease of temperature of the hot feed solution and increase of temperature of the cold distillate water flowing in the module. It is known that membrane distillation is a process driven by temperature difference. The higher the heat loss, the less heat is available for water evaporation on the membrane surface. The evaporation efficiency (η) of the process is defined as the ratio of the heat which contributes to evaporation and the total heat exchanged between the feed and the distillate. Therefore, η can be expressed as

η ) NV∆HVS/V˙ feedcp∆Tf ) NV∆HVS/V˙ distillatecp∆Tp (6) Here, V˙ feed and V˙ distillate represent the volumetric flow rates of the feed and distillate respectively, ∆HV is the heat of vaporization of water having a specific heat cp, and ∆Tf and ∆Tp are the temperature drop of the feed and the increase of distillate temperature along the module length, respectively.

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The evaporation efficiencies at a brine inlet temperature of 85 °C in the seven rectangular cross-flow membrane modules are summarized in Table 2. Clearly modules MXFR 1 and MXFR 3 have very high evaporation efficiencies compared to the others; this came about due to the smaller conductive heat losses, suitable membrane module design, and very high water vapor fluxes. MXFR 3 having coated PP 150/330 fibers is potentially an excellent membrane module for DCMD with the hot brines. How to shorten the residence time of distillate in the tube side is one of the important ways to ensure that the DCMD system has a larger driving force. The PP 150/330 fibers in modules MXFR 1 and MXFR 3 have a large i.d. (330 µm) which allows these membranes to have high distillate flow rate in the tube side so as to effectively reduce the temperature increase of the cold distillate stream. Meanwhile, the large fiber bore allows the distillate flow pressure drop to remain low (Table 2). The pressure drop is also directly proportional to the length of flow path. Therefore, the pressure drop of the distillate through the tube side of the module 4 is much higher than that through module MXFR 6 since the PP 50/200 fibers in module 4 are almost 3 times longer than those in the module MXFR 6 as shown in Table 1. An important aspect of our research in DCMD is the capacity to predict the water vapor flux through a given membrane in a specific module under given operating conditions. This is a complex problem due to coupled heat and mass transfer in a composite membrane having a complex device design. A number of factors are crucial here. (1) There is a considerable amount of conductive heat loss. The higher the wall thickness and the higher the porosity, the lower is the conductive heat loss. (2) The water vapor flux is considerably affected by the fiber coating resistance. (3) Unlike heat exchanger tubings, which are rigid, the hollow fibers used here are flexible. (4) The distillate-side heat transfer conditions are also important. Vacuum membrane distillation (VMD) studies can eliminate the first and the fourth factors since there will be no conductive heat loss and distillate-side resistance (and the corresponding distillate-side axial temperature profile). Our next publication will address this analysis by studying VMD in the same modules. This is expected to isolate the brine side heat transfer coefficient for the flexible hollow fiber-based cross-flow modules as well as the membrane mass transfer resistances which will facilitate analysis of the performance of the DCMD modules. 4. Concluding Remarks Seven rectangular cross-flow hollow fiber modules and one parallel flow module were investigated for their DCMD performances with hot brine in cross flow over the outside surface of the hollow fibers and cold distillate flowing in the hollow fiber bore. A variety of plasmapolymerized microporous silicone-fluoropolymer coatings were applied on the o.d. of the porous hydrophobic PP hollow fibers having three different i.d.s and two different wall thicknesses. The DCMD study varied the brine temperature between 60 and 90 °C as the velocities of the hot brine and the distillate were varied over a wide range. The hollow fibers in module MXFR 3, having the largest wall thickness (150 µm) and fiber i.d. (330 µm), as well as the most permeable composite of microporous coating-on-pore as determined via N2 permeation, yielded

at high brine velocities the highest water vapor flux in DCMD, namely, in the range of 41-79 kg/m2‚h for hot brine feed between 85 and 90 °C. These fibers in module MXFR 3 had the lowest conductive heat loss and the highest heat transfer efficiency among all modules/ fibers. The hot brine Reynolds number range needed for such performances was not very high (around 30-75). A considerable increase in the brine side heat transfer coefficient due to cross flow helped to achieve the high water vapor flux by reducing the temperature polarization in the brine. Larger fiber bore accompanied by high distillate flow rate is essential to keeping the distillate temperature rise to a low value and achieving a high water vapor flux. An 120-hr-long continuous DCMD run with 85 °C brine, in addition to further tests for an additional 280 h, did not show any evidence of salt leakage into the distillate. However, prefiltration would be useful to prevent fouling of the hollow fiber surfaces by dust/ particles, etc., and control any consequent flux reduction. The water vapor flux results obtained are highly encouraging and are as good as, if not better than, the best reverse osmosis membrane performances. The results further suggest that scale-up of this DCMD membrane and device for similar flow conditions would be highly desirable. Acknowledgment We acknowledge funding for this research from the Desalination and Water Purification Research and Development Program (Contract 01-FC-81-0737) Bureau of Reclamation, Denver, Colorado. We appreciate the assistance of Dr. S. Majumdar and Ms. Y. Wu. Notation cp ) liquid heat capacity d ) characteristic dimension di ) fiber inside diameter (i.d.) do ) fiber outside diameter (o.d.) DCMD ) direct contact membrane distillation DI water ) deionized water Hg ) mercury i.d. ) internal diameter kg ) kilogram L ) fiber length MD ) membrane distillation n ) number of fibers in a membrane module NV ) mass flux of water vapor across the membrane o.d. ) outside diameter pfm ) water vapor partial pressure at hot brine-membrane interface ppm ) water vapor partial pressure at cold distillatemembrane interface P1 ) atmospheric pressure QN2 ) permeability coefficient of N2 permeation through the membrane of effective thickness δM QN2/δM ) N2 permeance Re ) Reynolds number Red ) diameter-based Reynolds number RO ) reverse osmosis S ) inside membrane area () nπdiL) STP ) T0 ) 273.15 K, P0 ) 760 mmHg T ) temperature T1 ) room temperature Tf ) bulk temperature of feed

Ind. Eng. Chem. Res., Vol. 43, No. 17, 2004 5309 Tfm ) interface temperature on the surface of membrane in feed side Tp ) bulk temperature of distillate Tpm ) temperature of the distillate-membrane interface on the distillate side v ) velocity V ) volume of water transferred from feed to distillate V˙ 1 ) volumetric flow rate of gas through the membrane during measurement at room temperature V˙ feed ) volumetric flow rate of feed V˙ distillate ) volume flow rate of distillate VMD ) vacuum membrane distillation F ) density µ ) dynamic viscosity (absolute viscosity) η ) membrane heat transfer efficiency ∆HV ) heat of vaporization of water ∆Pbr ) breakthrough pressure for hollow fiber pore ∆PN2 ) nitrogen pressure difference across membrane corrected to STP ∆Tf ) temperature drop of feed along the module length ∆TP ) temperature increase of distillate along the module length δM ) effective thickness of membrane

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Received for review December 16, 2003 Revised manuscript received May 4, 2004 Accepted May 13, 2004 IE030871S