Recovery of Phenol from Aqueous Streams in Hollow Fiber Modules

Apr 2, 2002 - Recovery of Phenol from Aqueous Streams in Hollow Fiber Modules. W. Cichy and J. Szymanowski *. Institute of Chemical Technology and Eng...
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Environ. Sci. Technol. 2002, 36, 2088-2093

in water than TBP. Owing to all of these factors, TBP is not considered at present as a perspective reagent for the recovery of organic acidic substances, including phenol.

Recovery of Phenol from Aqueous Streams in Hollow Fiber Modules W. CICHY AND J. SZYMANOWSKI* Institute of Chemical Technology and Engineering, Poznan University of Technology, Pl. M. Sklodowskiej-Curie 2, 60-965 Poznan, Poland

A setup with two parallel hollow-fiber modules was used to study the recovery of phenol from aqueous solutions. Cyanex 923, Amberlite LA-2, and trioctylamine (TOA) in aliphatic kerosene were used as carriers. A solution of 0.2 M NaOH was used for stripping. It was found that each of the studied carriers permitted the effective removal of phenol. Cyanex 923 showed the best performance, removing phenol in the shortest time and giving the highest fluxes and the highest mass-transfer coefficients. The maximum fluxes of phenol entering the receiving phase changed in the following ratio: Cyanex 923/Amberlite LA-2/TOA ) 3.5/ 1.5/1. The mass-transfer coefficient in the extraction step changed in the same order: 34/5.2/1. The masstransfer coefficients of the stripping step were 2-4 orders lower than in the extraction step and were comparable for each carrier: Cyanex 923/Amberlite LA-2/TOA ) 1.1/0.7/ 1. Using Cyanex 923, only 5 min were needed to recover 99% of the pollutant from the aqueous stream, containing 0.5-2 g L-1 phenol.

Introduction Phenol and its derivatives are commonly encountered in aqueous effluents from various manufacturing processes (1). Solvent extraction, with such organic solvents as benzene, heptane, toluene, methyl isobutyl ketone, diisopropyl ether, isopropyl ether, isopropyl acetate, and so forth, was used for phenol recovery (2). Phenol is a weak acid (pKa ) 9.98) and can be efficiently extracted with various basic and solvating reagents, including a variety of alkylamines (e.g., Amberlite LA-2, Alamine 336) (3), tributyl phosphate (TBP), trialkylphosphine oxides (4-7), and trialkylphosphine sulfide (8, 9). The reaction of phenol with solvating reagents occurs according to the following equation (7):

PhOHw + So ) PhOH‚So

(1)

The equation is also valid for separation of phenol with neutral amines (i.e., when the pH of the aqueous feed is high enough to avoid protonation of amines). TBP and trioctylphosphine oxide (TOPO) are the most popular solvating reagents, used for many years in the processing of spent nuclear fuel and recovery of uranium(V), respectively. TOPO is a stronger extractant as compared with TBP (also in the extraction of organic acids and phenol). TOPO is a stable reagent which, in contrast to TBP, does not undergo hydrolysis either in contact with aqueous acid or alkaline solutions. Moreover, TOPO shows lower solubility * Corresponding author fax: [email protected]. 2088

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TOPO is an individual compound with high purity. Moreover, being a solid, it must be dissolved in an appropriate solvent. The use of polar solvents, including aromatics, is preferred. Using them, high distribution ratios are achieved, but it violates the rules of green chemistry and environmental protection. New commercial technical trialkylphosphine oxides (Cyanex 921, Cyanex 923, and Cyanex 925) have been produced by Cytec Co. Their compositions were determined recently by us (10, 11). Cyanex 921 is a high-purity reagent and contains approximately 99% TOPO. Cyanex 923 and Cyanex 925 are mixtures of several alkylphosphine oxides. As mixtures, these extractants are more soluble in organic diluents, including aliphatic kerosene. Moreover, Cyanex 923 is a liquid and can be also used without any diluent. However, some phase disengagement problems occur in solvent extraction experiments. Therefore, the use of aliphatic kerosene is advantageous, especially in membrane experiments where phase viscosity has an important role to play. The production of Cyanex 925 has been stopped. Thus, Cyanex 923 is the best extractant from the group considered, and it extracts much more phenol than does TOPO (5). Cyanex 471X is another extractant with high purity. It contains tri(isobutyl)phosphine sulfide as the active substance. It is solid and must be used upon dilution. Having a sulfur atom instead of oxygen, Cyanex 471X is a less stable and weaker extractant than Cyanex 923. There are important drawbacks to classic dispersive extraction, for instance, third phase formation, emulsification, dissolution of reagent and diluent in the aqueous phase, and incidental leakage. As a result, solvent extraction may cause significant contamination of aqueous streams with spent chemicals. These problems can be eliminated using membrane processes. Moreover, such processes seem to be natural successors of traditional extraction processes, likely to develop quickly in the very near future (12). Hollow fiber modules were used by several authors to recover toxic substances, including phenol, from various aqueous effluents (13-18). Methyl isobutyl ketone, isopropyl acetate, and methyl tert-butyl ether were used as carriers for phenol recovery. The use of alkylamines and trialkylphosphine oxides as phenol carriers was not considered. Only Garea et al. (19) have used Cyanex 923 for phenol extraction with a supported liquid membrane. The recovery of phenol in an extraction-stripping process carried out in two parallel hollow-fiber modules has not been described earlier. However, the use of a sequence of two hollow-fiber modules has already been reported for lactate separation (20, 21). A comprehensive review of the applications and fundamentals of membrane processes was published recently by Sastre et al. (12). The aim of this work was to study the recovery of phenol from aqueous solutions in a membrane extraction-stripping process using three different carriers (trialkylamine, dialkylamine, and trialkylphosphine oxide) and to select the most suitable reagent for practical applications. Trioctylamine (TOA), Amberlite LA-2, and Cyanex 923 were selected as the best representatives of extractants from the groups of tertiary amines, secondary amines, and trialkylphosphine oxides, respectively. 10.1021/es010910w CCC: $22.00

 2002 American Chemical Society Published on Web 04/02/2002

FIGURE 1. Hollow fiber setup (F, M, R: feed, membrane, and stripping reservoirs, each of 0.5 L volume): (R) flowmeter; (HF) hollow-fiber module; (P) pump; (PG) pressure gauge; (SV) sampling valve; (V) valve; (W1, W2) reservoirs for waste aqueous phases at the first start of the process.

Experimental Section The following reagents were used as carriers: TOA (Fluka, Germany), Amberlite LA-2 (Fluka, Germany), and Cyanex 923 (Cytec Co.). Amberlite LA-2 contained N-dodecyl-N(1,1,3,3,5,5-hexamethylhexylamine) as the active substance, while Cyanex 923 contained various trialkyl (C6, C8) phosphine oxides as the active substance with their content above 92% (10). Aliphatic kerosene (Slovnaft, Slovakia), containing 99.80% aliphatic hydrocarbons (n-C10, 7.18%; n-C11, 32.39%; n-C12, 33.11%; n-C13, 26.84; and n-C14 and higher, 0.28%) was used as a solvent. Phenol, pure (POCh, Poland), was additionally distilled prior to the extraction experiments. The stock solutions in deionized water were stabilized with 0.1 mol of Na2SO4/1 g L-1 of phenol. The acidities of the aqueous solutions were adjusted with H2SO4 (3 and 0.2 M) and NaOH (2 and 0.2 M). Dispersive extraction was carried out in a typical manner at a room temperature using various phase ratios by volume. The phases were mechanically shaken during 3 h. The hollow-fiber setup as presented in Figure 1 had two Liqui-Cel 2.5 in. × 8 in. modules (HF1 and HF2), two peristaltic pumps (P1 and P2) (Masterflex I/P) one gear pump (ColeParmer, Vernon Hills, IL), a complete pressure controller (PWBlogg, Paul Wegener, Germany) with five sensors (PG1PG5) and three rotameters (R1-R3). Each phase was recycled. In the first hollow-fiber module HF1, phenol was extracted to the organic phase that was pumped to the second hollowfiber module, and there phenol was stripped with 0.05-0.2 M NaOH. The feed (pH 4) contained 0.5-2 g L-1 phenol and 0.05-0.2 M Na2SO4. The organic phase contained 0.4 M of the carrier, which was dissolved in aliphatic kerosene. Both aqueous phases flowed through the tubes of the modules. The organic phase flowed on the shell side. To start the process for the first time, the organic phase was pumped from the vessel M through the hollow-fiber HF2 and HF1 (always upstream in the modules). This helped wet the fiber pores and fill them with the organic phase. Next, the flow of the membrane phase was stopped, and deionized water was pumped through fiber tubes, upstream again, and collected in reservoirs W1 and W2 (valves V5 and V8 were closed and valves V6, V7, V9, and V10 were opened). The pressure was monitored with five sensors PG1-PG5, and the overpressure on the tube side was about 40 kPa to prevent any leakage of the membrane phase. When the organic phase, which leaked at the beginning of the process, was removed from the tube side (usually after a few minutes),

FIGURE 2. Distribution ratios of phenol (equilibrium pH range of 2-8 for Cyanex 923, 4-8 for TOA, and 6-8 for Amberlite LA-2; carrier concentration, 0.4 M). valves V6, V7, V9, and V10 were closed and valves V5 and V8 were opened. Thus, the countercurrent flows were established. The actual receiving phase was then introduced to the reservoir R and pumped through the tubes of the modules HF2. The water was pushed out from the module and from the pipeline and removed from the system. A 3-fold excess of the receiving phase was used, and the circuit was closed (i.e., the receiving phase started to be recycled). The actual feed solution was then pumped through the module HF2 in the same manner as was described for the module HF2. The organic phase was again pumped, and the extractionstripping process began. Samples (2 cm3 each) of the aqueous phases were taken at appropriate intervals. The content of phenol in the aqueous phases was determined spectrophotometrically. A Unicam 8625 spectrophotometer (U.K.) was used. The samples were diluted with 0.1 M NaOH to adjust the absorbance, and the content of ionic species (phenolate) formed was determined at 270 nm. The concentrations were calculated from calibration data (absorption coefficient amounted to 12.25 dm3 g-1 cm-1; standard deviation SD ) 0.0142). The organic phase was regenerated with 0.2 M NaOH aqueous solutions before each new experiment until no phenol was detected in the aqueous phase. After that, deionized water was used to remove sodium hydroxide. When experiments with a new organic phase were planned, all the three phases were removed, and the setup was washed with water and 2-propanol and dried with warm air.

Results and Discussion The distribution ratios of phenol defined by the ratio of phenol concentration in the organic and aqueous phases required to estimate the mass-transfer coefficients and determined in extraction experiments, as given in Figure 2, showed that the physical extraction of phenol by kerosene is small (D ) 0.12). Cyanex 923 (operating at pH 2-8) has a much higher distribution ratio than Amberlite LA-2 (operating at pH 6-8), which, in turn, has a higher distribution ratio than TOA (at pH 4-8), which, in turn, has a much higher distribution ratio than kerosene alone. The different ranges of pH resulted from the different basicity of the amine carriers (pKa ) 3.5 and 5.3 for TOA and Amberlite LA-2, respectively) and their protonation in the acidic environment. The distribution ratio did not depend on phenol concentration for the amine carriers but decreased with phenol concentration for Cyanex 923. In each case considered, the distribution ratio sharply VOL. 36, NO. 9, 2002 / ENVIRONMENTAL SCIENCE & TECHNOLOGY

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FIGURE 3. Phenol concentration in the aqueous feed (empty symbols) and receiving phase (full symbols) for various carriers: (F) 1 g L-1 of phenol, pHFo ) 6; (M) 0.4 M of TOA, Amberlite LA-2, and Cyanex 923; (R) 0.2 M NaOH; UF ) UR ) 2.4 L min-1 and UM ) 0.8 L min-1; (points) experimental data: (O, b) Cyanex 923, (4, 2) Amberlite LA-2, (0, 9) TOA; (solid line) estimation according to the model. decreased at pH values above 8, and no extraction was observed for an aqueous phase containing 0.05-0.2 M NaOH. Extraction experiments indicated that Cyanex 923 was the strongest extractant among the reagents considered. Cyanex 923 was also a stronger extractant than Cyanex 471X. Schlosser et al. (8) reported that the distribution ratio of phenol extracted by Cyanex 471X was equal only to about 1.6 under similar extraction conditions (i.e., using the same kerosene as a diluent and a similar molar concentration). Thus, the obtained results fully confirmed the best extraction properties of Cyanex 923 in respect to phenol. The best extraction properties of Cyanex 923 with respect to phenol were thus fully validated by the results obtained. Stripping of phenol was not studied in the published works (3-9). However, we observed disengagement problems in stripping with aqueous NaOH solutions. Emulsions were formed particularly when Cyanex 923 was used. Therefore, the reagent was not suitable for classic dispersive extraction but could be successfully used in the membrane process with limited phase contact. Some representative kinetic data, given in Figures 3 and 4 indicate that extraction of phenol with Cyanex 923 is very effective. Within 3 min, 97% of the phenol was removed from the feed. However, about 30 min was necessary to transfer phenol to the receiving phase. No leakage occurred when the pressure on the aqueous side was 40-100 kPa higher than on the kerosene side. These results show how phenol extraction can be carried out using Cyanex 923 without emulsification. High recoveries of phenol were also possible with both amine reagents, but the transfer of phenol was slower, and about 60 min was needed to transfer phenol to the receiving phase. The kinetic data given in Figure 3 agreed well with the model proposed by Coelhoso and co-workers (20, 21) for the extraction-stripping of lactate (solid lines in Figure 3). The following two typical mass-transfer relationships were formulated:

-VF

dcF ) KF AF(cF - cF*) dt

dcR VR ) KRAR(cR* - cR) dt

(2)

(3)

where the equilibrium concentration of phenol was calculated 2090

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FIGURE 4. Phenol concentration in the organic phase for various carriers: (F) 1 g L-1 of phenol, pHFo ) 6; (M) 0.4 M of TOA, Amberlite LA-2, and Cyanex 923; (R) 0.2 M NaOH; UF ) UR ) 2.4 L min-1 and UM ) 0.8 L min-1; (points) experimental data: (b) Cyanex 923; (2) Amberlite LA-2, (9) TOA. from the distribution ratio

cF* ) cM/DF

(4)

cR* ) cM/DR

(5)

The concentration of phenol in the organic phase was calculated from the mass balance

cM )

VFocFo - VF cF - VRcR VM

(6)

As is presented in Figure 1, the distribution ratio was constant in both systems containing amines (DF ) 1.35 for TOA and DF ) 2.20 for Amberlite LA-2), and log D changed in the system containing Cyanex 923, approximately according to the linear relationship

log D ) 2.07 - 0.513cF

(7)

Because of sampling, the volumes of recycled aqueous phases decreased from the initial value VF0 ) 0.570 L and VR ) 0.480 L to about 0.540 and 0.450 L, respectively (depending on the number of samples taken for analysis). Taking into account that the volume changes did not exceed 7%, the average values were taken into consideration. Actually, the concentration of phenol should also be the function of the hollow-fiber length. However, that change could be totally neglected for amines used as the carriers because the flows of phases were high (0.8-4.0 L min-1) and the tube side volume of hollow-fiber modules was only 0.145 L. Thus, the average contact time was in the range 2-10 s and only 3 s for the typical flow of 2.4 L min-1. As a result, only a small change of phenol concentration between the inlet and the outlet of the module was observed (0.02, 0.03, and 0.11 g L-1 for TOA, Amberlite LA-2, and Cyanex 923, respectively, equivalent to the change of 2%, 3%, and 11% at the start of the process with the highest gradient and flux). Thus, the existence of gradient concentration along the module could be totally neglected in systems containing amines, and one could not expect any important approval of the model for the system containing Cyanex 923. The change of the phenol concentration in reservoirs F and R could also be neglected because of small volumes of

TABLE 1. Fluxes of Phenol from Membrane to Receiving Aqueous Phasea carrier TOA

Amberlite LA-2

Cyanex 923

UF ) UR* L min-1

JR 105 mol m s-2

UM** L min-1

JR 105 mol m s-2

1.2 2.0 2.8 3.6 4.0 4.7 0.8 1.6 2.4 3.2 4.0 0.8 1.6 2.4 2.8 3.2

0.421 0.661 0.550 0.550 0.538 0.601 0.788 0.801 0.826 0.790 0.790 1.78 1.93 1.81 1.99 1.93

0.4 0.8 1.6 2.4

0.474 0.550 0.504 0.574

0.4 0.8 1.6 2.4 3.2 0.8 1.2 2.0 2.8

0.804 0.826 0.896 0.780 0.845 1.93 1.83 2.04 1.95

a (F) 1 g L-1 phenol, pH Fo ) 4; (M) 0.4 M solutions; (R) 0.2 M NaOH; (*) UM ) 0.8 L min-1; (**) UF ) UR ) 0.8 L min-1.

FIGURE 5. Initial flux of phenol entering the membrane: (F) 0.5-2 g L-1 of phenol, pHFo ) 4; (R) 0.2 M NaOH; (M) 0.4 M in kerosene; UF ) UR ) 2.4 L min-1, UM ) 0.8 L min-1. liquids (0.300 L) in comparison with the flows of recycled streams and changes of concentration in the hollow-fiber modules. As solving eqs 2 and 3, the mass-transfer coefficients were computed. The model (solid lines in Figure 3) fitted well the experimental data. The mass-transfer coefficients for the extraction step changed in the following order: TOA < Amberlite LA-2 < Cyanex 923, reaching the values 4.05 × 10-5, 2.10 × 10-4, and 1.38 × 10-3 m s-1, respectively. The highest mass-transfer coefficient in the stripping step was also obtained for Cyanex 923, but the differences between values obtained for different carriers were relatively small. The mass-transfer coefficients were 6.53 × 10-7, 4.25 × 10-7, and 7.36 × 10-7 m s-1 for TOA, Amberlite LA-2, and Cyanex 923, respectively. Thus, the stripping step proceeded more slowly than the extraction step, as the differences between the mass-transfer coefficients were significant (2-4 orders of magnitude). At the selected flow rates (UF ) UR ) 2.4 L min-1 and UM ) 0.8 L min-1), the results are consistent with the highest resistance observed in the pores of the fibers, as an increase in the flows of the aqueous phases up to 4.7 L min-1 and in the membrane phase up to 3.2 L min-1 did not cause any increase of either the mass-transfer coefficients or the fluxes of phenol.

FIGURE 6. Maximum flux of phenol entering the receiving phase: (F) 0.5-2 g L-1 of phenol, pHFo ) 4; (R) 0.2 M NaOH; (M) 0.4 M in kerosene; UF ) UR) 2.4 L min-1, UM ) 0.8 L min-1. Table 1 shows the effect of linear velocities of the feed and the receiving phase, when varied simultaneously at a constant linear velocity of the membrane phase, on the phenol flux to the receiving phase (columns 2 and 3), as well as the effect of the linear velocity of the membrane at the constant flows of both aqueous phases (columns 4 and 5). Actually, the results deviated only slightly from the average, showing that the flux amounted to 1.91 × 10-5, 0.815 × 10-5, and 0.542 × 10-5 mol m s-2 for Cyanex 923, Amberlite LA-2, and TOA, respectively, indicating the best performance for Cyanex 923. Important differences between the mass-transfer coefficients calculated for the extraction and stripping steps could not be explained by various resistances of diffusion in the pores of the fibers in the extraction and strip hollow-fiber modules, respectively. The contribution of the interfacial reaction had to be considered. It is now generally accepted that extraction-stripping processes, in which hydrophobic reagents such as those used in this work are involved, occur in the interfacial region with preadsorption of hydrophobic carriers (22, 23). The interfacial extraction step could be described by the following equations:

SF ) SAd

(8)

PhOHint + SAd ) PhOH‚SAd

(9)

PhOH‚SAd ) PhOH‚SM

(10)

with eq 9 as the slowest one. The concentration of the neutral reagent S in the adsorption layer was high with respect to the bulk concentration and could be determined from the interfacial tension isotherm (22-26). In the case considered, the interface seemed to be saturated with the surface excess equal to 1-3 × 10-6 mol m-2 (24-26). Such advantageous conditions did not exist at the strip (membrane-receiving phase) interface. The complex PhOH‚ S was not amphiphilic, and it did not accumulate at the interface. Thus, the interfacial concentration of the complex was low. Moreover, the liberated carrier was adsorbed at the strip interface, blocking it and forming additional interfacial resistance. The phenomenon was similar to that observed in extraction systems containing surfactants or amphiphilic impurities (22, 23). As a result, the stripping step was significantly slower than the extraction step. The adsorption of reagents and the mechanism of the interfacial reaction also explained well the observations of VOL. 36, NO. 9, 2002 / ENVIRONMENTAL SCIENCE & TECHNOLOGY

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fluxes were observed after 3-5 min (Figure 8). Thirty and 60 min was needed to decrease the fluxes near zero for Cyanex 923 and Amberlite LA-2 and TOA, respectively.

Acknowledgments The work was supported by the University Grant DBP No. 32/3371999-2001.

Notation

FIGURE 7. Flux of phenol entering the membrane: (F) 0.5-2 g L-1 of phenol, pHFo ) 4; (R) 0.2 M NaOH; (M) 0.4 M in kerosene; UF ) UR ) 2.4 L min-1, UM ) 0.8 L min-1; (b) Cyanex 923, (2) Amberlite LA-2, and (9) TOA.

A

interfacial area, m2

c

mass concentration, g dm-3

D

distribution ratio

J

flux density, mol m-2 s-1

K

mass-transfer coefficient, m s-1

U

flow rate, L min-1

V

volume, m3

t

time, min

Indices Ad

adsorption

F

feed phase

int

interfacial region (film)

lm

logarithmic mean

M

membrane phase

o

initial value

R

stripping solution, interface on the strip side M/R

*

equilibrium value

Literature Cited FIGURE 8. Flux of phenol entering the receiving phase: (F) 0.5-2 g L-1 of phenol, pHFo ) 4; (R) 0.2 M NaOH; (M) 0.4 M in kerosene; UF ) UR ) 2.4 L min-1, UM ) 0.8 L min-1; (b) Cyanex 923, (2) Amberlite LA-2, and (9) TOA. other authors concerning the recovery of various carboxylic acids (20, 21, 27). In each case, irrespective to the type of liquid membrane and the type of the carrier, the stripping step was slower than the extraction step. The statement was also valid for other membrane processes, including the bulk liquid membranes and the recovery of metal species from various waste streams (28-30). The initial flux of phenol entering the membrane and the maximum flux entering the receiving phase (i.e., observed for the maximum concentration of phenol in the membrane) was strongly dependent upon the concentration of phenol in the feed (Figures 5 and 6). Approximately linear relationships were observed. Only in the case of Cyanex 923 did JR deviate from the straight line for phenol above 1.2 g L-1, in agreement with the effect of phenol concentration on the distribution ratio (Figure 2). The fluxes of phenol entering the membrane decreased rapidly during 10 min (Figure 7). The drop was especially strong for Cyanex 923. As a result, the main quantities of phenol were transferred to the membrane in this period of time. After 10 min, only negligible amounts of phenol were further extracted with fluxes similar for each carrier. Phenol was observed to accumulate in the membrane. The driving force for stripping increased and the maximal 2092

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(1) Mulligan, T. J.; Fox, R. D. In Industrial Wastewater and Solid Waste Engineering; Cavaseno, V., Ed.; McGraw-Hill Publications Co.: New York, 1980; pp 173-190. (2) King, C. J.; Senetar, J. J. In Ion Exchange and Solvent Extraction; Marinsky, J. A., Marcus, Y., Eds.; Marcel Dekker: New York, 1988; Vol. 10, pp 35-61. (3) Cichy, W.; Schlosser, S.; Szymanowski, J. Proceedings of the 25th Conference SSCHI, on CD ROM with full texts, Jasna´, Slovakia, 1998; pp 1-9. (4) MacGlashan, J. D.; Bixby, J. L.; King, C. J. Solvent Extr. Ion Exch. 1985, 3, 1. (5) Watson, E. K.; Rickelton, W. A.; Robertson, A. J.; Brown, T. J. Solvent Extr. Ion Exch. 1988, 6, 207. (6) Urtiaga, A. M.; Ortiz, I. Sep. Sci. Technol. 1997, 32, 1157. (7) Drozdova, M. K.; Nikolaeva, I. V.; Torgov, V. G.; Russ, J. Phys. Chem. 1997, 71, 498. (8) Schlosser, S.; Rothova, I.; Frianova, H. J. Membr. Sci. 1993, 80, 99. (9) Schlosser, S.; Rothova´, I. Sep. Sci. Technol. 1994, 29, 765. (10) Dziwin ˜ ski, E.; Szymanowski, J. Solvent Extr. Ion Exch. 1998, 16, 1465. (11) Dziwin ˜ ski, E.; Szymanowski, J. Solvent Extr. Ion Exch. 1999, 17, 1515. (12) Sastre, A. M.; Kumar, A.; Shukla, J. P.; Singh, R. K. Sep. Purif. Methods 1998, 27, 213. (13) Sengupta, A.; Basu, R.; Sirkar, K. K. AIChE J. 1988, 34, 1698. (14) Prasad, R.; Sirkar, K. K. AIChE J. 1988, 34, 177. (15) Basu, R.; Prasad, R.; Sirkar, K. K. AIChE J. 1990, 36, 450. (16) Yun, C. H.; Prasad, R.; Sirkar, K. K. Ind. Eng. Chem. Res. 1992, 31, 1709. (17) Boyadzhiev, L.; Lazarova, Z. In Membrane Separation Technology. Principles and Applications; Noble R. D., Stern S. A. Eds.; Elsevier Science: Amsterdam, The Netherlands, 1995; pp 283352. (18) Nanoti, A.; Ganguly, S. K.; Goswani, A. N.; Rawat, B. S. Ind. Eng. Chem. Res. 1997, 36, 4369.

(19) Garea, A.; Urtiaga, A. M.; Ortiz, M. I.; Alonso, A. I.; Irabien, J. A. Chem. Eng. Commun. 1993, 120, 85. (20) Coelhoso, I. M.; Crespo, J. P. S. G.; Carrondo, M. J. T. J. Membr. Sci. 1997, 127, 141. (21) Coelhoso, I. M.; Silvestre, P.; Viegas, R. M. C.; Crespo, J. P. S. G.; Carrondo, M. J. T. J. Membr. Sci. 1997, 134, 19. (22) Szymanowski, J. Hydroxyoximes and Copper Hydrometallurgy; CRC Press: Boca Raton, FL, 1993; pp 1-304. (23) Szymanowski, J. Solvent Extr. Ion Exch. 2000, 18, 729. (24) Prochaska, K. Adv. Colloid Interface Sci. 2002, 95, 51. (25) Radzio, K.; Prochaska, K. J. J. Colloid Interface Sci. 2002, 233, 211. (26) Prochaska, K.; Walczak, M.; Staszak, K. J. Colloid Interface Sci. 2002, 248, 143.

(27) Schlosser, S.; Sabolova´, E. Chem. Pap. 1999, 53, 403. (28) Kirschling, P.; Nowak, K.; Miesiac, I.; Nitsch, W.; Szymanowski, J. Solvent Extr. Res. Dev., Jpn. 2001, 8, 135. (29) Alonso, A. I.; Urtiaga, A. M.; Zamacona, S.; Irabien, A.; Ortiz, I. J. Membr. Sci. 1997, 130, 193. (30) Boyadzhiev, L.; Dimitrov, K. Solvent Extr. Ion Exch. 1996, 14, 105.

Received for review May 1, 2001. Revised manuscript received February 13, 2002. Accepted February 20, 2002. ES010910W

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