Comparison of Liquid Membrane Processes for ... - ACS Publications

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Ind. Eng. Chem. Res. 2003, 42, 392-403

Comparison of Liquid Membrane Processes for Metal Separations: Permeability, Stability, and Selectivity X. J. Yang,*,† A. G. Fane,† and K. Soldenhoff‡ UNESCO Centre for Membrane Science and Technology, University of New South Wales, Sydney 2052, Australia, and Australian Nuclear Science and Technology Organization, PMB 1, Menai, NSW 2234, Australia

A comparison of bulk, supported, and contained liquid membranes (BLM, SLM, and CLM) for metal separations with respect to permeability, stability, and selectivity is presented. The overall mass-transfer coefficient of magnitude order of 10-6 and 10-7 m s-1 was typically found for SLM/ BLM and CLM, respectively. The SLM has the highest organic utilizing efficiency and the poorest membrane stability. Osmotic pressure across the membrane proves to be one of the major causes of SLM instability, and membrane pore elongation due to morphology changes reduces the SLM lifetime. The separation efficiency of liquid membranes can be enhanced by minimizing the organic inventory and maximizing the contact area between the aqueous phase and the organic membrane phase. However, this needs to be balanced against the improvement in stability of various liquid membrane configurations, which can be achieved by employing a significant organic inventory. The selectivity of liquid membranes mainly depends on the extractant used in the membrane phase and is similar to that of one-stage solvent extraction. The paper concludes with a technical comparison of the various liquid-membrane techniques. 1. Introduction Since its first industrial application in the nuclear fuel cycle process for uranium purification after the Second World War, solvent extraction (SX) has become a mature technology, with wide-ranging applications in the mining/metallurgy, petroleum, pharmaceutical, and food industries. The advantages of SX include high throughput, ease of automatic operation and of scaleup, and high purification. However, the most commonly used contacting equipment, such as pulsed columns, mixer settlers, and centrifugal contactors, suffer from many disadvantages. These include the need for dispersion and coalescence, problems of emulsification, flooding, and loading limits in continuous countercurrent devices, the need for density differences between the phases, phase disengagement difficulties, high solvent losses, and large solvent inventories. On the other hand, SX is an equilibrium-based separation process, where the extraction and stripping are two separate steps and therefore the separation that can be ultimately achieved is limited by the conditions of equilibrium. Liquid-membrane (LM) processes combine extraction and stripping into one single stage and thus have nonequilibrium mass-transfer characteristics where the separation is not limited by the conditions of equilibrium. Of the LM processes, the emulsion LMs (ELMs) and supported LMs (SLMs) had been extensively studied since the 1970’s. An ELM pilot plant for zinc recovery from wastewaters was first reported in the late 1980’s.1 However, the ELM industrialization has undergone

difficulties because of the complication of the process and the membrane instability (membrane swelling and breakage). The SLM technology eliminates the majority of the problems encountered in conventional SX, and the simplicity of the SLM processes has led to active research in the late 1980’s and early 1990’s. A SLM pilot plant for recovery of copper was reported in 1998.2 However, the permeability of copper decreased by 60% after a period of 500 h of operation because of stability problems. The instability of the SLMs has become the major obstacle for industrialization of the processes.3 On the other hand, the nature of the instability mechanisms is under dispute.4 This has impacted the improvement of SLM stability. Since the early 1990’s, new LM configurations such as hollow fiber contained LM (HFCLM) and electrostatic pseudo-LM (ESPLIM) have been developed for improvement of the membrane stability.5,6 Although each configuration has been claimed to avoid the instability problem, industrialization of the so-called new LM technologies is not yet extensively practiced. Comparisons of the various LM processes are needed in order to clarify the advantages and disadvantages of the different LM configurations. In this study a brief overview of LM processes is presented, followed by an experimental comparison of the bulk LM (BLM), supported LM (SLM), and contained LM (CLM) with regard to the parameters of permeability, stability, and selectivity. The aim of this study is to provide guidelines for the development of LM technology in practical applications. 2. Principle and Applications of LM Processes

* To whom correspondence should be addressed. Present address: Environmental Chemistry Section, Environment Protection Authority, P.O. Box 29, Lidcombe, NSW 1825, Australia. Tel: 61 2 9995 5045. Fax: 61 2 9646 2755. E-mail: [email protected]. † University of New South Wales. ‡ Australian Nuclear Science and Technology Organization.

BLMs usually consist of an aqueous feed and stripping phase separated by a bulk organic LM phase. In its simplest form, this technique can be carried out in a U-tube configuration in the laboratory (Figure 1A, left). However, the BLM principle also applies to other LM processes, such as the CLM and ESPLIM (see below).

10.1021/ie011044z CCC: $25.00 © 2003 American Chemical Society Published on Web 12/20/2002

Ind. Eng. Chem. Res., Vol. 42, No. 2, 2003 393

Figure 1. Schematics of (A) a BLM, (B) an ELM, (C) a SLM, and (D) a CLM.

ELMs may be in water-organic-water (W/O/W) or organic-water-organic (O/W/O) form, with the middle phase essentially the LM phase (Figure 1B). This middle phase is a thin “shell” around the droplets where the continuous phase and the core of the droplet are the same phase (i.e., water with an organic shell). For membrane preparation, the first step commonly involves the formation of a stable emulsion of the stripping solution in the solvent phase, with the aid of a surfactant (normally 5-10%). The stable emulsion is then dispersed in a third continuous phase (the feed solution to be treated) to form a double emulsion. Extraction and separation take place by the transport of the solute(s) from the continuous outer feed phase through the LM phase to the inner stripping phase. Since the late 1980’s, industrialization of the ELM has been carried out. However, large-scale industrial applications of the ELM technology have been limited. The major problems of the ELMs are emulsion swelling and membrane rupture as well as complicated operation requiring emulsion formation and breaking. SLMs offer a simpler configuration and process compared to ELMs. A microporous hydrophobic membrane is impregnated with an organic solvent (which is typically used in SX) and then sandwiched between an aqueous feed and strip solution (Figure 1C). The microporous membrane for the SLM can be flat-sheet or hollow fibers (HFSLM). In the HFSLM the fiber wall is impregnated with the LM phase, the feed solution flows through the lumen of fibers, and the strip solution flows through the shell side of the fibers or vice versa. The

major disadvantage of the SLM is a lack of long-term membrane stability due to gradual loss of the membrane phase out of the membrane pores, which have been attributed to a variety of factors, such as lateral shear forces, static pressure differential and osmotic pressure across the membrane, and progressive wetting.7 These factors have been recognized by most researchers. However, some authors have disputed that the osmotic pressure improves SLM stability.8,9 The osmotic pressure effect is a topic of SLM stability in this study. Recently, Klaassen and Jansen10 reported an “emulsion pertraction” process that employed a considerable amount of the organic membrane phase dispersed with aqueous droplets of the strip phase for heavy-metal removal. Ho and Poddar11 reported a similar process (called SLM with strip dispersion) for removal and recovery of chromium from wastewaters. In these processes, the stability due to the loss of the organic phase out of the membrane pores is no longer a problem. However, it was observed that the transport takes place under the carrier saturation mechanism, i.e., equilibrium condition. In other words, the process no longer has nonequilibrium mass-transfer characteristics. In addition, these processes require emulsion formation and phase separation. CLMs utilize two sets of membranes, between which the membrane phase is contained (Figure 1D). Similarly to the SLM, both flat-sheet and hollow fiber membranes can be used. When hollow fibers are used in a single module, the technique is known as the HFCLM.12-14 In the HFCLM, the strip and feed solutions are passed

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(ii) Extraction: forward chemical reaction of eq 1

ke ) Kex

(

[HR]2

[H+]f2

)

-1

(4a)

or

ke ) DaKd/δfb

(4b)

(iii) Cu-LIX complex diffusion in the LM phase Figure 2. Schematic representation of the transport processes in an SLM.

km ) Dm/δm

(5)

In the case of SLM, eq 4b becomes through the lumen of two separate sets of fibers, with the solvent membrane phase being stationary on the shell side; this uses the principle of the BLM. This process can also be carried out in two separate hollow fiber modules with the membrane phase circulating between the two modules.15 This configuration is called the hollow fiber membrane contactor (HFMC). Other LMs that have been claimed as new types, such as flowing LMs16 (similar in principle to the CLM), creeping film pertraction,17 and rotating film pertraction18,19 (similar in principle to the BLM), are hybrid techniques with some elements of SX. ESPLIMs were obtained by a combination of electrostatic, SX, and LM techniques. The principle has been described previously.20,21 A pilot plant for yttrium recovery from wastewaters was reported.22

In this study, copper transport was selected for studies of permeability through LMs in most of the experiments because the SX chemistry of Cu2+ by an extractant of LIX agent has been well established and that in the LMs is exactly the same as that in the SX

Cu

2+

+

+ 2HR S CuR2 + 2H Kex )

[CuR2][H+]2 [Cu2+][HR]2

Kd )

[Cu] [Cu2+]

(1)

(2a)

(2b)

where the overbar denotes the species in the membrane (organic) phase, HR stands for the extractant LIX agent, Kex is the extraction equilibrium constant, and Kd is the distribution coefficient. However, the stripping process (the reverse reaction of eq 1) takes place simultaneously with the extraction process at the other side of the membrane and, therefore, the extraction equilibrium expressed by eq 2 no longer exists as in the conventional SX process. Figure 2 depicts schematically the Cu transport processes in the LM process, which have the transport coefficients as follows: (i) Cu2+ diffusion in the aqueous feed/membrane boundary layer

kfb ) Da/δfb

(3)

(6)

(iv) Stripping: reverse chemical reaction of eq 1. (v) Diffusion in the aqueous strip/membrane boundary layer. (vi) Back-diffusion of the regenerated LIX agent in the LM phase. The copper flux was obtained by

J)-

Vf dCf A dt

(7)

Integration of eq 7 leads to the measurement of the overall mass-transfer coefficient, koverall

ln

3. Transport Process

Dm δm τ

km )

Ct kA t )C0 V

(8)

To compare the performance of different LM processes, a new parameter, the membrane-utilizing efficiency coefficient (MUE), is defined by the transported amount by unit volume of the membrane phase at unit time:

MUE )

Vs dCs Vo dt

(9)

A limited number of experiments were also performed using the BLM, SLM, and HFSLM on the zirconium/ hafnium separation system in order to examine the selectivity of LMs. The separation factor for two metal ions is given by

SF )

(MZr/MHf)product (MZr/MHf)feed

(10)

4. Experimental Section 4.1. Reagents and Membranes. The reagents included copper(II) sulfate pentahydrate CuSO4‚5H2O (Ajax, Australia; analytical grade), LIX 984N (Henkel; 50% oxime in kerosene, a 1:1 mixture of 5-nonylsalicylaldoxime and 2-hydroxy-5-nonylacetophenone oxime), kerosene (Aldrich; d 0.800, n 1.4420, bp 175-325 °C), sulfuric acid (Ajax, Australia; 98%), tri-n-octylamine (TNOA; 95%, Fluka), 2-ethyl-1-hexanol (Aldrich; 99%), sodium hydroxide (Ajax, Australia), and potassium chromate (Ajax, Australia). TNOA, tributyl phosphate (TBP), kerosene, 2-ethyl-1-hexanol, cyclohexane, zirconium, and hafnium standard solutions (1 mg L-1) were obtained from Aldrich Chemical Co., Inc. All other chemicals used were of analytical grade, and all of the

Ind. Eng. Chem. Res., Vol. 42, No. 2, 2003 395 Table 1. Reported Parameters of Membranes Used membrane

membrane type

pore size (µm)

thickness (µm)

porosity (%)

tortuosity

material

Celgard 2500 Celgard 2502 Accurel PP-2E Accurel Q3/2a

flat sheet flat sheet flat sheet hollow fiber

0.05 × 0.19 0.075 × 0.25 0.2 0.2

25 50 150 200 (wall)

45 45 69 69

1.80-3.0 1.80 1.37

polypropylene polypropylene polypropylene polypropylene

a

The fiber has an inside diameter of 600 µm (mean value) and an outside diameter of 1000 µm.

chemicals were used as received. Water with a resistivity of 18 MΩ‚cm purified using a Milli-Q (Millipore) system was used throughout. All reagents were used as received. The membrane supports for the SLM were flat-sheettype Celgard 2500 and 2502 and Accurel PP-2E and hollow fiber type Accurel Q3/2. Table 1 lists the reported properties of these membranes. 4.2. LM Apparatus and Operational Procedure. For permeability studies, the feed phase was an aqueous CuSO4 solution prepared from CuSO4‚5H2O at a copper concentration of 2 g L-1 and pH of 2.5 (the pH was adjusted by adding 2 M H2SO4). The strip phase was a 2 M H2SO4 solution. The LM phase was 10% LIX 984N in kerosene. The SX of copper with LIX 984N was previously studied,23 and the system was applied to the studies of the stability of the SLM.24,25 For selectivity studies, the feed phase was an aqueous mixture of zirconium and hafnium (each at 25 mg L-1) in 8 M HCl and the strip phase was a 1 M HCl solution. The LM phase was 25% (v/v) TBP in kerosene or 0.5 M TNOA and 1 M 2-ethyl-1-hexanol in kerosene (2-ethyl-1hexanol was used as a phase modifier for Zr and Hf extraction). TBP and TNOA have been extensively used as extractants for Zr and Hf extraction.26-28 A systematic study for the transport of Zr and Hf through hollow fiber supported LMs was also previously reported.29 BLM. The inner dimension of the transport cell was 150 mm long × 60 mm wide × 95 mm deep. A barrier with 60 mm height divided the cell into two compartments. Both feed and stripping solutions were 220 mL and were stirred by magnetic stirrers at 110 rpm at the bottom of the cell. The dimension of the magnetic bar with Telflon coating was 7 mm in o.d. × 35 mm long. The volume of the membrane phase was 220 mL. Both interfaces (feed/membrane and membrane/stripping) were at the same level and were separated by a barrier above the interface of 8 mm. SLM. The flat-sheet SLM cell consisted of halves of a cylindrical chamber (36 mm in i.d. × 75 mm long) separated by the membrane. The effective volume of each chamber was 75 mL, and the contact area of the membrane was 1.02 × 10-3 m2. The SLM was prepared by impregnating a membrane disk (50 mm diameter) with the organic solution (10% v/v LIX 984N for copper extraction and 0.5 M TNOA or 25% TBP for Zr and Hf extraction) for at least 15 min (the membrane became uniformly transparent). Then, the membrane was taken out of the organic solution, and the excess organic liquid attached to the surface of the membrane was removed gently with a tissue. The weight ratios of the LM phase over the membranes Celgard 2500 and 2502 and Accurel PP-2E were 0.8-0.85, 0.75-0.8, and 2.4-2.6, respectively. The prepared membrane was then placed in the SLM cell. The feed and strip phases were then simultaneously introduced into the respective chambers. A magnetic stirrer agitated the aqueous solutions in each chamber at 450 rpm. It was found that the masstransfer flux was independent of the stirring speed

Figure 3. Flowsheet of HFSLM for removal and recovery of copper from wastewater.

above 300 rpm. Therefore, a stirring speed of 450 rpm was set as the standard stirring condition in this study. CLM. A cell with a volume of 41 mL was sandwiched between the halves of the cylindrical chamber (Figure 1D). The feed and strip solutions were introduced. Then, the middle cell was filled with 40 mL of 10% LIX 984Nkerosene. An overhead stirrer was placed in the middle cell to stir the organic phase at 300 rpm. The aqueous feed and strip phases were stirred with magnetic bars at 450 rpm. HFSLM. (a) Copper extraction. A glass tube with an i.d. of 30 mm (o.d. 35 mm) and a length of 410 mm was used for housing the fibers. The sealing material at the two ends of the glass tube was epoxy resin. The number of fibers was 420, and the effective length in the module was 340 mm. The packing density was 46%, and the specific area was 1500 m2 m-3. The total volume of 10% LIX 984N-kerosene absorbed in the fibers was 65 mL. A feed solution (pH 2.5, Cu 350 ppm) was pumped through the lumen of the fibers at a flow rate of 1 L h-1. The strip solution (200 mL of 2 M sulfuric acid) was recycled through the shell side of the fibers at a flow rate of 2 L h-1. The flowsheet of the process is shown in Figure 3. (b) Zr/Hf transport. A glass tube of 18 mm o.d. with an effective length of 300 mm was used. The packing density of the module was 46%. The volume of 0.5 M TNOA absorbed in the fibers was 22 mL. The feed phase (200 mL of 8.5 M HCl containing 25 ppm of Zr and Hf) was recirculated through the lumen of the fibers at a flow rate of 2 L h-1. The stripping phase (200 mL of 1 M HCl) was recirculated through the shell of the fibers at the same flow rate. 4.3. Analytical Methods. The copper concentration was determined by ethylenediaminetetraacetic acid titration or atomic adsorption spectrometry (Varian AA275), and Zr and Hf concentrations were measured with inductively coupled plasma atomic emission spectroscopy. Samples (0.5 mL) were taken from the feed and strip chambers for metal concentration measurement. 4.4. Characterization of Membranes (SEM). The membrane samples were cut into small pieces (5 mm long × 2.5 mm wide; for imaging the cross section of the membranes, the samples were freeze fractured in

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Table 2. Mass-Transfer Flux, Coefficient, and MUE in BLM, CLM, and SLM in Copper Transport

LM

feed/ organic ratio

initial feed flux (mol m-2 s-1)

initial strip flux (mol m-2 s-1)

overall flux (mol m-2 s-1)

estimated km (m s-1)

measured koverall (m s-1)

MUE (mol m-3 s-1)

BLM CLM Celgard 2500 SLM Celgard 2500 SLM Celgard 2502 SLM Accurel PP-2E

1 1.9 3000 1200 720

5.5 × 10-5 1.4 × 10-5 4.9 × 10-5 3.5 × 10-5 1.7 × 10-5

1.2 × 10-5 7.3 × 10-7 4.8 × 10-5 2.3 × 10-5 8.8 × 10-6

1.0 × 10-5 1.3 × 10-6 1.5 × 10-5 1.7 × 10-5 7.4 × 10-6

2 × 10-6 1 × 10-6 2 × 10-6 1 × 10-6 7 × 10-7

2 × 10-4 3.5 × 10-5 0.6 0.4 0.03

3.0 × 10-6

5 × 10-7

1.8 × 10-6 2.2 × 10-7 1.8 × 10-6 2.3 × 10-6 3.9 × 10-7 (0-40 h) 1.3 × 10-6 (45-80 h) 6.6 × 10-7

HFSLMa a

538

0.01

The feed solution for the HFSLM was an electroplating wastewater containing 350 mg L-1 Cu.

Figure 4. Copper transport kinetics through (A) BLM, (B) CLM, and (C) SLM. Feed: 2 g L-1 Cu2+, pH 2.5. LM: 10% LIX 984N in kerosene. Strip: 2 M H2SO4.

liquid nitrogen) and glued to specimen holders. The specimens were then coated with about a 2 nm thick chromium film using an MED 010 Turbo mini deposition system (Balzers Union Ltd.) and then imaged using a field emission scanning electron microscope (FESEM; S-900 Hitachi). The pore size was measured using an Image Processing Toolkit (IPTK) inside Photoshop 5.0. 5. Results and Discussion 5.1. Permeability. The transport kinetics of the BLM, CLM, and SLM are illustrated in Figure 4. In the BLM and CLM, the increase in the copper concentration in the stripping phase does not match the decrease in the copper concentration in the feed phase, while the increase in the stripping phase is equal to the decrease in the feed phase in the SLM process. Consequently, the copper accumulation in the membrane phase of the BLM and CLM is significant, and that in the SLM is negligible. Figure 4 also shows that copper was trans-

ported from the feed solution to the strip solution (2 M H2SO4) against its concentration gradient (in the SLM process, 50% of Cu was transported to the stripping phase at an operating time of 8 h and 95% transported at 40 h; Figure 4C). This is an important characteristic of the LM process. The mass-transfer fluxes and coefficients are given in Table 2. The initial strip flux was calculated as 4.8 × 10-5 mol m-2 s-1 for SLM-Celgard 2500 and was 4 times lower for the BLM (at 1.2 × 10-5) and 66 times lower for the CLM (at 7.3 × 10-7). The CLM system has the lowest transfer rate, and the BLM and SLM systems have rates similar to those of the BLM. The SLM has a major advantage in that the organic inventory is very small for a similar transfer rate. The membrane masstransfer coefficient, km, was calculated from eq 6. For this calculation, a value of 2 × 10-10 m2 s-1 predicted by the Wilke and Chang equation,30 which is supported by diffusion data of Danesi et al.31 and Alguacil and Sastre,32 was taken for calculating the mass-transfer coefficient across the membrane [the tortuosities of the Celgard 2500 and Accurel membranes were taken from the literature (Table 1)33]. The calculated values of the membrane transfer coefficient (km) were 2 × 10-6 and 7 × 10-7 m s-1 for the Celgard 2500 and Accurel PP2E, respectively (Table 2). The mass-transfer coefficient across the feed boundary layer (kfb) and the interfacial mass-transfer coefficient (ke) were estimated to be 6 × 10-6 and 8 × 10-4 m s-1 (Da ) 6 ×10-10 m2 s-1;34 δfb ) 100 µm;35 Kd ) 13036). The Kex for Cu extraction by LIX 860 was measured as 5.4 × 104.37 LIX 984N is a mixture of LIX 860 and LIX 84, both of which have been extensively used as an industrial extractant for copper extraction.36 Therefore, the mass-transfer coefficient in the Celgard 2500 and Accurel PP-2E SLMs was on the order of ke > kfb > km ≈ koverall. It may be concluded that the overall mass-transfer process in the SLM is controlled by diffusion in the microporous membrane. However, some authors reported that the aqueous boundary layer diffusion and membrane diffusion were the simultaneous controlling factors in a SLM of 100 mg L-1 Cu2+||GVHP-LIX 54||0.9 M H2SO4.32 Komatsu et al.38 and Danesi et al.31 reported that the rates of Cu2+ reaction with LIX reagents are several orders of magnitude higher than the permeation rates. Lee et al.39 found that the mass-transfer resistance was negligible on the strip side of the membrane in a SLM system (835 mg L-1 Cu2+||Celgard 2500-4.5% LIX 65N||1 N H2SO4) and that membrane diffusion was the chief resistance to mass transfer. O’Hara and Bohrer40 found that diffusion in the aqueous boundary layer and interfacial complexation and decomplexations were not controlling the rate of transport of the Cu2+-LIX 54-SLM system. Valenzuela et al.41 considered that diffusion of Cu-LIX 860 through the microporous membrane was the rate-

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Figure 5. Plot of -ln(Ct/C0) vs operation time in the SLM process: (b) Celgard 2500; (9) Accruel PP-2E.

controlling step in HFSLM for copper extraction. These results support our observation. However, in the case of CLM (where the same aqueous phases, LM phase, and microporous membranes as those in the SLM were employed), the mass-transfer resistance was shifted from membrane diffusion to the interfacial reaction. The mass-transfer coefficient of the CLM-Celgard 2500 was 1 order of magnitude lower than the SLM-Celgard 2500, and the initial feed and strip fluxes of the SLM were 3.5 and 65 times those of the CLM, respectively (Table 2). Because the hydrodynamic conditions of the feed and strip phases were the same, the aqueous boundary layer resistances would be the same. Therefore, the interfacial stripping reaction and the extraction reaction were the limiting steps of the CLM process. However, Guha et al.42 reported that both the feed aqueous boundary layer and the interfacial complexation reaction on the feed side dominate copper transport in a HFCLM with LIX 84-n-heptane for feed copper concentrations of 90-500 mg L-1. The comparison between the CLM and SLM presented above would suggest that the mass-transfer rates of the HFCLM and HFMC would be lower than that of the HFSLM for the same hollow fiber membranes and aqueous-membrane contact area. This has been experimentally observed by Valenzuela et al.43 Guha et al.42 observed that the transfer rate could be increased by increasing the membrane area in the HFCLM. Obviously, this will increase the capital costs, and a compromise between the transfer rate and the capital costs should be considered in practical applications. The significantly lower transfer rate of the Accurel membranes compared with the Celgard membranes was due to its greater thickness and the greater amount of the organic phase. As a result, the Accurel membrane has a larger membrane transfer resistance than the Celgard membranes. It is interesting that the slope of -ln(Ct/C0) against time t for the first period of 40 h is significantly lower than that for the following period of 40 h for the Accurel membrane (Figure 5). This indicates that the overall mass-transfer coefficient, koverall, in the initial stage (first 40 h) was smaller than that in the later stage for Accurel SLMs (Table 2). This can be explained by the decrease in the LM thickness due to the loss of the organic phase out of the membrane pores. This further supports the membrane diffusion control in SLM systems. Copper extraction from electroplating wastewaters containing 350 mg L-1 Cu, 50 mg L-1 Cr, and 90 mg L-1 Zn was carried out using a HFSLM.44 The membrane performance (i.e., the mass-transfer rate and the MUE) of the HFSLM (Accurel Q3/2) was

comparable to that of flat-sheet SLM (Accurel PP-2E) in spite of the different experimental conditions (Table 2). The results demonstrate the ability of the LM technique for extraction and concentration from dilute streams. The major advantage of the SLM is the highest MUE (Table 2). Under the same conditions of hydrodynamic, aqueous feed and strip, and organic membrane phase compositions, the MUE of the SLM systems was on the order of Celgard 2500 > Celgard 2502 > Accurel PP-2E (Table 2) despite a reverse effective contact area order (Accurel PP-2E > Celgard 2502 > Celgard 2500) in accordance with the membrane pore size and porosity. This suggests that the thinner the membrane (i.e., less organic inventory), the higher transport efficiency. The significantly low MUE values of the BLM and CLM compared to the SLM indicate that the BLM and CLM approach the equilibrium status. The MUE (0.03 mol m-3 s-1) of the flat-sheet Accurel SLM was comparable to that (0.01 mol m-3 s-1) of HFSLM, although hydrodynamic conditions and membranes were different. The comparison of the CLM and SLM using flat Celgard and Accurel membranes in this study was carried out at the same conditions of hydrodynamic, aqueous feed and strip, and organic membrane phase compositions. This illustrates that the separation efficiency can be enhanced by using membranes as thin as possible (reducing the membrane transfer resistance) and minimizing the inventory of the organic phase. In other words, the nonequilibrium condition can be achieved by minimizing the membrane thickness (i.e., minimizing the organic inventory) and maximizing the contact area between the aqueous phase and the organic membrane phase. Guha et al.42 achieved overall mass-transfer coefficients of around (2-4) × 10-6 m s-1 in a HFCLM with a membrane thickness of 25 µm. The recently reported SLM with strip dispersion11 employing hollow fibers with a thickness of 30 µm had an overall mass-transfer coefficient of 3.2 × 10-6 m s-1 for chromium transport. The values of the overall mass-transfer coefficients of Guha et al.’s system and Ho and Poddar’s system were comparable to those obtained in this study. However, Ho and Poddar11 reported that their system was at the equilibrium condition. Therefore, the value of the masstransfer coefficient seems not able to judge if the process is operated under the nonequilibrium or equilibrium status. 5.2. Stability. The major problem of the SLM is the instability and limited membrane lifetime,4 and most investigators established the stability of SLMs empirically over the period of a couple of hours to months,36 depending on the organic solvents and membranes. In this study, it was found that the initial copper flux of the Celgard 2500 SLM decreased by 80% after 80 h of operation. It was observed that the copper started diffusing back to the feed phase from the strip phase at 110 and 350 h of continuous operation for the Celgard 2500 and Accurel PP-2E, respectively. Those critical times can be regarded as the lifetime of the corresponding SLM.25 During the course of SLM transport, the LM phase is gradually removed from the pores of the membrane. This phenomenon has been regarded as direct evidence of SLM instability. Chemical reactions are usually involved in SLMs, and this may cause gradual changes to physical properties, such as interfacial tension and viscosity, etc., which influence the LM phase. Therefore, the stability of the SLM process is one

398


of the most difficult parameters to predict from a theoretical basis. The mechanisms of SLM instability have been the subject of many studies, and Kemperman et al.4 presented a comprehensive review. The mechanisms reported include (1) effect of LM phase solubility in the adjacent aqueous phases (feed and strip), (2) effect of the membrane support, (3) effect of transmembrane pressure, (4) effect of osmotic pressure, (5) shear-induced emulsion of the LM phase, (6) wetting of the membrane pores by aqueous phases, and (7) contamination and pore-blockage. Mechanisms 1 and 2 cover the effect of the physical properties and structure of the LM phase and the membrane support. Mechanisms 3-5 belong to the effect of operating conditions. Mechanisms 6 and 7 might reflect the results of chemical reactions. Mechanisms 1-3 have been well evident and recognized in the SLM process. However, mechanisms 4-7 have been disputed.4 Among these, the most disagreement is mechanism 4, the effect of osmotic pressure. Fabiani et al.45 and Danesi et al.7 reported that the osmotic pressure decreased the stability, whereas Neplenbroek et al.,8,9 Kemperman et al.,4 and Zha et al.46 claimed that the osmotic pressure improved the stability. Osmotic pressure as a possible mechanism of SLM degradation was first observed independently by Danesi et al.7 and Fabiani et al.45 This effect was also considered by several others as one of the causes of SLM instability.47-52 Danesi et al.7 reported that the presence of an osmotic pressure gradient over the SLM could lead to the displacement of the organic membrane phase from the membrane pores by water flux. The larger the osmotic pressure gradient, the greater the quantities of water transported through the membrane. The movement of water first occurred through organic-filled pores and gradually displaced the organic phase within the membrane pores. Eventually, when the membrane pores became devoid of the organic phase completely, the facilitated transport through the SLM disappeared. Fabiani et al.45 measured water flows through SLMs consisting of the Celgard 2500 membranes impregnated with a trilaurylamine hydrochlorate-diethylbenzene solution and LiCl solutions and concluded that the organic phase was dragged out of the membrane pores as a consequence of the water volume flux across the membrane due to the osmotic pressure difference. However, the osmotic mechanism was challenged by Neplenbroek et al.8,9 They reported that the SLM stability was increased when an osmotic pressure difference across the membrane was created by an increase in the salt concentration in the stripping phase. Neplenbroek et al. did observe water transport across SLMs in the presence of an osmotic pressure gradient but claimed that this osmotic water flow was a consequence of the SLM instability rather than the cause of the SLM instability. The discrepancy might be due to the different organic phases and experimental conditions used, in particular, mutual solubility of the organic phase and water. When the water solubility in the organic phase is very low, contradictory conclusions might be drawn because of the different experimental conditions and the stability measurement methods. In fact, no organic solvent and water are absolutely mutually insoluble. It has been well recognized that the osmotic pressure gradient reduces the stability of the SLM when the water has a

Figure 6. Water transport through untreated microporous membranes under the driving force of an osmotic pressure gradient: (O) water||Accurel PP-2E||2 M H2SO4 + 1 M KCl; (b) water||Accurel PP-2E||2 M H2SO4; (9) water||Celgard 2500||2 M H2SO4. Stirring speed: 450 rpm.

Figure 7. Sulfuric acid concentration changes as a function of time in the two chambers of the SLM (the membranes were impregnated with kerosene): (A) Celgard 2500; (B) Accurel PP2E; (O) water chamber; (b) 2 M H2SO4 chamber. Stirring speed: 450 rpm.

high solubility in the LM phase. Deblay et al.48 observed that the SLM was reasonably stable for the organic solvent with solubility of less than 15 g L-1. When the solubility increased from 15 to 40 g L-1, the lifetime of SLMs decreased significantly, and for higher than 40 g L-1, the lifetime approached zero. To further investigate the effect of osmotic pressure, we measured osmotic water flows across the original hydrophobic membranes (untreated) and organic-treated membranes. Water transport through the untreated hydrophobic membranes under osmotic pressure is shown in Figure 6. It can be seen that the higher the osmotic pressure (i.e., the higher the salt concentration), the greater the water transport rate. In these experiments no SO42- transport from the sulfuric acid chamber to the water chamber was observed. The water transport is driven by the water vapor difference and is known as “osmotic distillation”. When the membrane was impregnated with the organic phase (as SLM), the situation was completely different. No net water flow from the water chamber to the acid chamber was observed, and sulfuric acid permeation from the acid chamber to the water chamber was observed. Eventually, the concentrations of sulfuric acid in the two chambers became equal (Figure 7). In contrast, Fabiani et al.45 observed that water flows into the chamber of the concentrated salt solution in a SLM system and that the water transport rates reached values similar to those of the untreated Celgard 2500 membrane when complete loss of the organic phase occurred. Fabiani et

Ind. Eng. Chem. Res., Vol. 42, No. 2, 2003 399 Table 3. Effect of the Osmotic Pressure on the Lifetime of SLM Systems (Membrane Celgard 2500)

a

system no.

chamber 1

organic phase

chamber 2

osmotic pressure gradient (atm)a

lifetime (h)

A B C D

H2O H2O 1.8 M NaCl 5 M NaCl

kerosene kerosene kerosene kerosene

2 M H2SO4 2 M H2SO4 + 2.5 M Na2SO4 2 M H2SO4 2 M H2SO4

84 166 0 149

230 112 298 220

Calculated from the equation Π ) (RT/V1)M1 × 10-3mφn. Table 4. Measured Surface Porosity and Pore Size of the Celgard 2500 before Use and after Use pore breadth (µm) pore length (µm) surface porosity (%) min max ave min max ave manufacture value before use after use

Figure 8. H+ concentration changes in water chambers as a function of time. The SLM systems are defined in Table 4.

al.’s results can be explained if the organic phase loss leaves empty membrane pores. However, it seems more plausible that the organic phase is replaced by the aqueous phase while it is gradually removed from the membrane pores. Consequently, those pores were filled with the aqueous phase and began channeling between the two aqueous phases, isolating the membrane. This mechanism has been explored by the use of impedance spectroscopy in a proceeding study53 and explains well the results presented in Figure 7. This channeling phenomenon is a sign of the SLM breakdown, and the time when this occurs can be regarded as the lifetime of the SLM. The lifetime of the SLM under different osmotic pressures was experimentally measured (Table 3). It clearly demonstrates that the SLM lifetime is longer when subject to a smaller osmotic pressure difference. The H+ permeation as a function of time in the SLMs is shown in Figure 8. The inflection points reflect the membrane leakage/channeling, i.e., the complete loss of organic solvent out of some of the membrane pores leading to direct connection of the two aqueous phases across the membrane. In all systems A-D (Table 3), no net water permeation was observed, whereas SO42permeation from chamber 2 into chamber 1 was observed, even in the case of system D, which has higher osmotic pressure in chamber 1. This can be explained by a solution-diffusion mechanism in terms of the effect of water flux “drag” on the organic phase within the membrane pores under the influence of osmotic pressure. If the water was completely insoluble in the organic phase, no water drag would occur even if a high osmotic pressure difference existed across the membrane. Our observation supports Danesi et al.’s mechanism. Significant changes to the membrane morphology were observed after the SLM transport process (Table 4 and Figure 9). The average pore size of the fresh membrane was in excellent agreement with the values provided by the manufacturer, indicating a good accuracy of the measurement by a FESEM. After a 650-h

0.05 11.5 10.8

0.19

0.012 0.15 0.051 0.024 0.53 0.19 0.009 0.29 0.051 0.015 0.49 0.13

SLM experiment, the pores became dilated and more elliptical.25 The pore breadth increased by 100%, and there was a slight decrease in the pore length after the SLM transport process. Therefore, for long-term SLM operation, the membrane morphology changes are potentially important factors affecting the lifetime and stability of the SLMs. Danesi et al.7 has provided guidelines for maximizing the lifetime of SLMs, and this study has provided additional insights. On the basis of previous studies,7,24,25,46,53 the augmented guidelines are summarized as follows: (i) microporous membranes having small pore radii and uniform pore size, less pore interconnectivity, high hydrophobicity, good resistance to organic solvent and acid, and considerable thickness, (ii) an organic LM phase having low solubility in water and low water solubility and high organic/water interfacial tension, (iii) minimization of the surface shear forces and transmembrane pressure gradients, (iv) minimization of the electrolyte concentration difference between feed and strip solutions separated by the SLM to reduce the osmotic pressure gradient, and (v) the instability of SLM processes involves a complex interaction of a number of factors, including surface shear forces, the Marangoni effect, osmotic pressure, membrane materials and morphology, Bernard instabilities, composition of membrane and aqueous phases, and membrane preparation protocol.24,25,54 5.3. Selectivity. It should be noted that the membrane itself has no selectivity and the high selectivity of the LM comes from the extractant used as a carrier in the LM phase. The SX is an equilibrium-based separation process, while LM is a nonequilibrium process. Therefore, it is worth comparing the selectivity of SX and LMs. Izatt et al.55 reported a comparison of the BLM, ELM, and SLM in the transport of alkalimetal ions and Zn, Cd, and Hg using macrocyclic compounds (crown ethers) and observed identical selectivities for different configurations when the same carrier was used. In a previous study, Zr-Hf pair separation was employed as the model system for the HFSLM because of the extreme chemical similarity of Zr and Hf. The comparison of Zr and Hf transport through the BLM, flat-sheet SLM, and HFSLM is shown in Figure 10, and the mass-transfer fluxes and separation factors of the LM systems are summarized in Table 5 in comparison with a single-stage SX. It is concluded that the selectivities of the LMs (BLM and SLM) can be similar to that of a single-stage SX. However, we have previously observed that the separation of two

400


Figure 9. Distribution of pore breadth and length of the Celgard 2500 membrane: (A and B) original membrane; (C and D) after 650 h of SLM transport. Table 5. Mass Fluxes and Separation Factors (SF) of Zr and Hf Obtained by the BLM, CLM, SLM, and HFSLM (Operation Time 60 h) and SX (Contact Time 10 min) overall mass flux (mol m-2 s-1)

Figure 10. Transport of Zr and Hf through the BLM and SLMs: (0) feed Zr, (9) feed Hf, (O) strip Zr, (b) strip Hf. LM phase: 0.5 M TNOA. (A) BLM: feed, 220 mL; strip, 220 mL; LM, 80 mL. (B) SLM: Accurel PP-2E; feed, 75 mL; strip, 75 mL. (C) HFSLM: feed, 200 mL; strip, 200 mL.

components/solutes, having similar thermodynamic extraction equilibrium but having quite different kinetics, can be improved in the ESPLIM.21 This is because the contact time is very short, which favors components with fast extraction kinetics. The separation of Zr and Hf by the HFSLM has been investigated in detail with respect to the effects of HCl

system

Zr

Hf

SF

BLM (TNOA) SLM (TNOA-Celgard 2500) SLM (TNOA-Accurel PP-2E) HFSLM (TNOA-Accurel PP2E) SX (TNOA, phase ratio ) 1:1) SLM (TBP-Accurel PP-2E) SX (TBP, phase ratio ) 1:1)

4.0 × 10-8 8.8 × 10-9 2.1 × 10-8 1.9 × 10-6

2.4 × 10-9 2.0 × 10-10 9.4 × 10-10 1.6 × 10-7

8.3 × 10-8

1.7 × 10-8

8.1 13.2 12.3 11.6 13.4 3.1 3.9

concentration, fiber length, and flow rate.29 It was shown that HFSLM for Zr and Hf separation can be operated with negligible impact of concentration polarization. The selectivity of SLM could be improved to some extent by decreasing the boundary film and membrane resistance on the basis of the thermodynamic selectivity of the extractant. 5.4. Qualitative Comparison of LM Process Configurations. It has been realized that quantitative comparison of the various LM processes becomes difficult because no quantitative criteria can be applied to the different LM processes. However, from a qualitative view of point, the qualitative comparison may be of general guidance. As for the permeability, it is impractical to set a value of the mass-transfer coefficient to compare the efficiency of different LM processes, in particular when the membrane area and thickness are different. However, the MUE proposed in this study is comparable for different LM processes. The SLM has the highest MUE. As for the selectivity, the microporous membrane itself has no selectivity contribution and the process selectivity is provided by the extractant. However, extraction kinetics may make a considerable contribution to the selectivity of the ESPLIM and ELM. The stability problem of SLMs is due to the loss of the organic phase. The BLM and CLM have no such stability problems because of the increased inventory of the organic phase, but this is at the cost of a decrease in the mass-transfer rate. Basu and Sirkar12 reported

Ind. Eng. Chem. Res., Vol. 42, No. 2, 2003 401 Table 6. Summary of Major Characteristics of Various LM Systems process

advantages

BLM ELM

useful for laboratory studies large-scale piloting experience available very high transfer rates

SLM

HFCLM

simpler process than ELM very low volumes of carrier required high interfacial area-to-volume ratio (500-5000 m2/m3) in spiral wound or HF modules high specific area in HF modules

HFC

commercially available HF modules easy to scale-up

ESPLIM

very versatile hybrid system that can be specifically designed to suit different extraction and stripping requirements highest concnetration factors high throughputs

that membrane life and stability problems encountered in SLMs were eliminated in the HFCLMs. However, pressure control across the membranes in the HFCLM operation can be more difficult than that in the HFSLM operation. In the long-term operation of HFCLM, organic phase leakage was observed.12 The ESPLIM technology does not suffer from stability problems because of the loss of the organic phase and can also show a high transport efficiency. However, the ESPLIM suffers from a problem of durability of electrodes immersed in the organic phase in a high electrostatic field. Duration of the electrodes was about 3-5 months in continuous operations. After this period, the electrodes must be replaced; otherwise, the electric spark occurs because of the damage/breakage of the insulation polyethylene layer. From the results of the LM techniques in this study and review of other LM techniques reported in the literature, we summarize in Table 6 the advantages and disadvantages of the various types of LMs. 6. Conclusions When the BLM, CLM, and SLM are compared, it is found that the SLM has the highest mass-transfer rate at the lowest organic inventory but suffers from instability and a limited membrane lifetime. Mass-transfer control in the SLM is shifted from feed boundary layer resistance for thin membranes to membrane diffusion for thicker membranes. The osmotic pressure and changes to the membrane pore size can play a very important role in the SLM stability during the SLM operation. The selectivity of the LM is similar to that of a single stage of SX when the same extractant is used, but it could be enhanced by means of extraction kinetics. It appears that wide application of the LM is promising until a new configuration that has relatively less complexity and ease of operation is developed.

disadvantages low fluxes due to very low interfacial area-to-volume ratio relatively complex process problem with emulsion swelling and membrane rupture which limits the level of concentration that can be achieved scale-up not straightforward stability of the LM still a problem low mass-transfer rates no large-scale long-term pilot studies available relatively complex HF module manufacture, not commercially available slightly higher solvent inventory when compared to SLM lower transport rates than SLM due to extra membrane resistance pilot-plant studies required operation pressure control difficulties slightly higher solvent inventory when compared to SLM and HFCLM lower transport rates than SLM due to extra membrane resistance operation pressure control difficulties complex cell to manufacture no commercial prototype available need for electrode maintenance

List of Symbols δ ) thickness (µm) ) porosity (%) τ ) tortuosity A ) membrane area (m2) C ) concentration (mol L-1, g L-1) D ) diffusion coefficient (m2 s-1) HR ) extractant J ) flux (mol m-2 s-1) k ) mass-transfer coefficient (m s-1) Kd ) distribution coefficient Kex ) extraction equilibrium constant M ) metal Q ) flow rate (L h-1) SF ) separation factor t ) time (s) V ) volume (L) y ) metal concentration in organic phase (mol L-1) Subscripts a ) aqueous phase e ) extraction f ) feed fb ) boundary layer of feed and membrane m ) membrane o ) organic phase r ) raffinate s ) stripping phase

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Received for review December 30, 2001 Revised manuscript received October 21, 2002 Accepted November 5, 2002 IE011044Z

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