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Recovery of sulfur aroma compounds using membrane-based solvent extraction
Journal of Membrane Science 187 (2001) 239–253
Recovery of sulfur aroma compounds using membrane-based solvent extraction F.X. Pierre, I. Souchon∗ , M. Marin UMR Génie et Microbiologie des Procédés Alimentaires, Institut National de la Recherche Agronomique INRA-INA PG, F-78850 Thiverval Grignon, France Received 3 November 2000; received in revised form 15 January 2001; accepted 17 January 2001
Abstract This work focuses on a non-destructive process for recovering valuable aromatic fractions from the food industry’s odorous wastewaters. Non-dispersive solvent extraction of three sulfur aroma compounds, dimethyldisulfide, dimethyltrisulfide and S-methyl thiobutanoate, was carried out from very diluted aqueous solutions representing real effluent. The mass transfer from water to n-hexane was studied using a cross-flow designed hollow fiber membrane contactor. A preliminary study showed high affinity of solutes for n-hexane, with constant partition coefficients at infinite dilution between water and hexane in a 90–560 range. The influence of tube and shell side hydrodynamics on mass transfer was studied, with the aqueous phase on the tube side, and the organic phase on the shell side. The diffusion of solutes from the bulk aqueous phase to the aqueous–organic interface controlled the separation and contributed, under the conditions tested, to more than 97% of the overall mass transfer resistance. A resistance-in-series model overestimated overall mass transfer coefficients. The main explanation is the inaccuracy of the Lévêque correlation used at low Reynolds numbers. The choice of a correlation for predicting mass transfers in the solvent phase did not affect the estimation, since the corresponding mass transfer resistance was negligible. Mass transfer fluxes obtained experimentally by membrane-based solvent extraction were greater for the three aroma compounds than those obtained by pervaporation. © 2001 Elsevier Science B.V. All rights reserved. Keywords: Membrane contactors; Liquid–liquid extraction; Sulfur aroma compounds; Modeling; Pervaporation
1. Introduction Food manufacturers often generate odorous effluents. Recent environmental constraints have led the food industry to focus on this problem with particular attention. Two kinds of approaches can be taken in order to treat odorous effluents. The first one consists of a process for destroying the odor, such as biofiltration, chemical reactions or incineration. ∗ Corresponding author. Tel.: +33-1-30-81-54-86; fax: +33-1-30-81-55-97. E-mail address: [email protected] (I. Souchon).
Since the molecules involved in odorous pollution are also present in numerous food products as flavoring compounds, it would be interesting to consider the non-destructive treatment of odorous effluent. In this second approach, techniques such as adsorption, absorption or membrane processes can be used. The objectives would in this case be to deodorize the effluent and recover a valuable “natural” aromatic fraction in order to compensate for the cost of the deodorization treatment. Sulfur compounds are the main odorous compounds responsible for olfactory pollution, in particular in food manufactures [1]. These sulfur compounds are
0376-7388/01/$ – see front matter © 2001 Elsevier Science B.V. All rights reserved. PII: S 0 3 7 6 - 7 3 8 8 ( 0 1 ) 0 0 3 5 2 - 0
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Nomenclature a A C d dh D DMDS DMTS e Gz k K LTU m ˙ MTB n ppb ppm P Q Re Sc Sh t v v¯ V
area per volume (m2 m−3 ) membrane area in the module (m2 ) concentration (kg m−3 ) diameter (m) hydraulic diameter (m) diffusion coefficient (m2 s−1 ) dimethyldisulfure dimethyltrisulfure membrane thickness (m) Graetz number local mass transfer coefficient (m s−1 ) overall mass transfer coefficient (m s−1 ) fiber length (m) length of a transfer unit (m) mass transfer flux (kg s−1 ) S-methyl thiobutanoate number of fibers in the bundle volumetric part per billion (1 ppb = 1 l m−3 ) volumetric part per million (1 ppm = 1 l l−1 ) liquid–liquid partition coefficient (kg m−3 /kg m−3 ) volumetric flow rate (m3 s−1 ) Reynolds number Schmidt number Sherwood number time (s) fluid velocity (m s−1 ) mean fluid velocity (m s−1 ) volume (m3 )
Greek letters β constant in Eq. (7) ε membrane porosity ε th geometric void fraction φ module packing density η dynamic viscosity (Pa s) µ kinematic viscosity (m2 s−1 ) ρ fluid density (kg m−3 ) τ membrane tortuosity Subscripts aq aqueous b in relative to the bundle inside diameter
b out i lm mb org shell t in t out tube
relative to the bundle outside diameter relative to solute i logarithmic mean membrane organic relative to the shell side of the module relative to the inside of the tubes relative to the outside of the tubes relative to the tube side of the module
Superscripts ∗ equilibrium value eq equilibrium ini initial
often the result of sulfate or sulfur organic compound reduction with anaerobic bacteria. But these compounds are also often naturally present in food products and can be considered as key aroma compounds, particularly in some cheeses and vegetables [2]. For example, cauliflower blanching water, which is generally spread, contains numerous sulfur compounds responsible for the odor [3]. In this former study, three very odorous sulfur compounds were identified: dimethyldisulfide, dimethyltrisulfide and S-methyl thiobutanoate. These molecules are involved in the aromatic quality of numerous French cheeses [4], and the development of an efficient process to recover them from wastewater processes would be very worthwhile. In the food industry, certain constraints must be taken into account when choosing the process to recover aroma compounds from liquid effluents. On the one hand, it is necessary to respect molecular integrity (no high temperature because of the sensitivity of aroma compounds temperature), and on the other hand to have an efficient process which is easy to install and maintain. The use of membrane contactors can be a technical solution taking all these criteria into account. Such membranes can be used to fix an interface between a liquid and either a gas or another liquid. This latter situation represents an alternative to liquid–liquid extraction referred to as membrane-based solvent extraction or non-dispersive solvent extraction. Non-dispersive extraction is performed by inserting a microporous membrane wall between the feed phase and the stripping phase. Using
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microporous membranes, one of the two fluids to be contacted wets the membrane pores and tends to leak on the other side. To prevent the mixing that would occur, slight overpressure must be applied to the other phase in order to stabilize the interface at the pore mouth. In spite of this, the technique remains a concentration-driven process. With respect to conventional liquid–liquid extraction, many advantages of the membrane contactors can be pointed out. There are no loading, flooding or emulsification problems; except for the pumping of the flowing phases, no agitation or moving part is needed. There is no need for density differences, leading to a greater choice of stripping solvents. A drawback occurs when using a membrane because it creates additional resistance that hinders diffusion from one phase to another, thus slowing the separation down. In most cases, the large surface area per volume offered by hollow fiber modules overcomes this disadvantage. Non-dispersive solvent extraction appears to be a competitive technique in different fields, including • chiral separation [5], • extraction of ethanol and organic acids produced by fermentation [6–8], • protein extraction [9], • pharmaceutical applications [10], • metal ion extraction [11], and • extraction of various pollutants from wastewaters [12,13]. The studies cited cover a wide range of partition coefficients (values from about 10−1 to 103 ), various solvents, membrane shapes and flow configurations. Only a few authors have studied the recovery of aroma compounds by membrane–solvent extraction. Fabre et al. [14] have studied the recovery of 2-phenylethylalcohol (rose-like aroma) produced by fermentation. These authors concluded that non-dispersive solvent extraction was a very promising method with respect to other techniques such as adsorption, pervaporation or simple liquid–liquid extraction. Spinnler et al. [15] have registered a patent for the recovery of gamma-decalactone (peach-like aroma) from fermentation broth by non-dispersive solvent extraction. Reed et al. [16] have presented an example of pilot scale membrane-based solvent extraction: the extraction of numerous organic pollutants was carried
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out from industrial wastewaters from two plants in Holland. However, industrial applications are scarce, due to the finite life of modules as well as the nature of potting adhesives that may be altered by certain organic solvents. The aim of this work was to apply a technology involving microporous membrane contactors to the recovery of sulfur compounds from very diluted aqueous solutions close to food manufacturer waste streams. In order to improve the separation process, a main objective was to improve the understanding of mass transfer in membrane contactors by identifying the controlling step. Thus, extraction experiments were carried out using a microporous hollow fiber membrane and dilute aqueous solutions representing industrial wastewaters. The influences of tube side and shell side hydrodynamics were studied and discussed, and a resistance-in-series model was used to predict mass transfer coefficients. The experimental performances obtained are discussed and compared to those obtained with another membrane technique, pervaporation.
2. Theoretical approach Membrane-based solvent extraction, like solvent extraction in traditional devices, is a concentration-driven operation. As a general rule, the mass flow rate of solute i (m ˙ i ) from one phase to another is defined by Eq. (1) m ˙ i = Ki A(Ci aq − Ci∗aq )
(1)
where K is the overall mass transfer coefficient based on the aqueous phase concentration, A the surface area of the membrane, Caq the concentration in the ∗ the hypothetical concentraaqueous phase, and Caq tion in an aqueous phase in equilibrium with the actual solvent phase. ∗ and C Caq org are related by the partition coefficient (P), that expresses the equilibrium state as eq
Ci org Ci org Pi = ∗ = eq Ci aq Ci aq
(2)
This equilibrium parameter is dependent upon temperature and concentration. Moreover, P is one of the main criteria for choosing the solvent. Indeed, the
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higher the partition coefficient, the more the equilibrium state is displaced towards the solvent phase, and the more complete the extraction. 2.1. The resistance in series model in hollow fibers Mass transfer is always described as a succession of diffusion steps, from the aqueous boundary layer, through the solvent filled membrane pores, to the solvent phase boundary layer. A resistance-in-series model is used to relate the overall mass transfer coefficient (K) to the local mass transfer coefficients (k). When the aqueous phase flows inside the tube of a hydrophobic hollow fiber membrane, and the organic phase flows outside the tube, the resistance-in-series model is expressed by Eq. (3) 1 1 dt in dt in = + + Ki aq Ki aq Pi ki mb dt lm Pi ki org dt out
(3)
where kaq and korg are the local mass transfer coefficients relative to the aqueous and organic boundary layers, respectively, and kmb the local mass transfer coefficient in the membrane; dt in and dt out are the internal and external fiber diameters, respectively, and dt lm their logarithmic mean value [17]. 2.2. Estimation of local mass transfer coefficients Several assumptions have to be made in order to use the simple expression for the membrane mass transfer coefficient estimation proposed by Kiani et al. [18]. • The membrane is symmetrical and its pores completely wetted by one phase. If the transmembrane pressure P is close to a critical pressure Pcr , partial displacement of the pore filling phase may occur, and the estimation of kmb becomes uncertain; • Diffusion of solutes in the pore-filling fluid must be unhindered. Beck and Schultz [19] suggest that diffusion can be considered to be unhindered if the pore diameter is about two orders of magnitude greater than the solute dimensions. Under these conditions, the mass transfer coefficient inside the membrane can be predicted [18] as ki mb =
Di ε τe
(4)
where Di is the diffusivity of the considered solute in the pore-filling fluid, ε and τ membrane porosity and tortuosity, respectively, and e is membrane thickness, replaced by (dt out −dt in )/2 in the case of hollow fibers. Gabelman and Hwang [20] have reported that the Lévêque solution is widely used in the literature to predict tube side mass transfer. The Lévêque solution (Eq. (5)) is a limiting case of the Graetz solution, applicable to a laminar flow in a tube when the Graetz number is high (Gz > 4), with: 0.33 0.33 0.33 dt in Shtube = 1.62 Retube Sc (5) where Re is the Reynolds number (Re = vd/η), Sc the Schmidt number (Sc = η/D), Sh the Sherwood number (Sc = kd/D) and dt in and the inside fiber diameter and length, respectively. Thus, knowing the Sherwood number, the resulting expression of the tube side mass transfer coefficient (ktube ) is 2 0.33 D v ktube = 1.62 (6) dt in A similar approach, based on the relationship between dimensionless groups, is used for estimating shell side mass transfer coefficients in hollow fiber modules. A summary of these equations, for various experimental systems, is given by Gabelman and Hwang [20]. All of them are in the form of Sh = aReb Sc1/3 , a and b often being fitted from experimental observations. Interestingly, the exponent parameter b over the Reynolds numbers varies from one equation to another, approximately from 1/3 to 1 depending on the geometry of the system studied. Actually, the packing density φ of the membrane module is often included in parameter(s) a and/or b, since it greatly influences the Sherwood number. As an example, Prasad and Sirkar [21] studied liquid–liquid extraction with various solutes and solvents, using a hollow fiber module. For pure counter-currents, they suggested predicting the shell side mass transfer coefficients using 0.33 dh Shshell = β(1 − φ)Re0.6 (7) Sc shell where β is 5.85 or 6.1, respectively for hydrophobic or hydrophilic hollow fibers; in the range of 0 < Re < 500 and 0.04 < φ < 0.4.
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Wickramasinghe et al. [22] studied mass transfers of oxygen from water into nitrogen using three different module shapes (axially wound, hollow fiber fabric and a vane module). These authors proposed an average correlation valid for all three designs: 0.33 Shshell = 0.8 Re0.47 shell Sc
(8)
For cross-flow hollow fiber modules, the velocity of the fluid flowing on the shell side, because of its radial component, is diameter-dependent. The determination of a local velocity implies firstly calculating a hydraulic diameter (dh ) and a geometric void fraction of the fiber bed (ε th ) for the shell side space, as suggested by Schöner et al. [23]. Assuming an ideal cross-flow mode, the mean velocity is obtained by integrating the local velocity from the axis to the wall of the module. Once the hydraulic diameter, geometric void fraction and mean velocity are known, it becomes possible to estimate a mean Reynolds number (Eq. (9)) and a Sherwood number (Eq. (10)), expressed as follows: Reshell =
Shshell =
ρ vd ¯ h 2ρQ (db out + db in )ln(db out /db in ) = µ πµ ndt out (9) kshell dh = kshell D
db2 out
− db2 in
− nd2tout
Dndt out
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responsible for the cauliflower blanching water odor. They are particularly hydrophobic and volatile species, with remarkably low limits of detection by the human nose, of the order of a part per billion (ppb). Their main physical properties are reported in Table 1. n-Hexane was chosen as the extracting solvent for several practical reasons. The sulfur aroma compounds are highly soluble in it, but it is almost insoluble in water. It is authorized in food processing and is widely used in aroma compound manufacturers. Knowing that its boiling point is well below those of sulfur compounds studied, the recovery of these solutes may be based on a simple distillation. Another solvent may be chosen considering many other criteria (direct availability, low price, low toxicity, high safety, etc.). As an example, the use of natural vegetable oils, which are usually good solvents for aroma compounds [24], would generate an aromatized oil possibly usable in a food process. Capillary grade (≥98%) n-hexane was supplied by Carlo Erba, and used as received. 3.2. Partition coefficients between water and n-hexane
(10)
3. Materials and methods 3.1. Chemicals The three solutes chosen for this work are those identified by Baudot et al. [3] as being mainly
A known volume of an organic solution of each solute, at different concentration levels, was carefully poured over a known volume of pure water into a separating funnel. After a 24-h settlement period at 25 ± 1◦ C, both phases were carefully separated and analyzed by gas chromatography. Mixing the two phases should be avoided in order to prevent any loss of hexane in the aqueous phase. It can be replaced by a contacting time long enough for the system to
Table 1 Properties of the sulfur compounds studied
Full name Formula Supplier Purity (%) Molecular weight (g mol−1 ) Boiling point at normal pressure (◦ C) Density at 20◦ C (kg l−1 ) Odor
DMDS
MTB
DMTS
Dimethyldisulfide H3 C–S–S–CH3 Aldrich 99 94 116 1.046 Onion, cauliflower
S-methyl butanoate H3 C–CH2 –CH2 –C(=O)–S–CH3 Aldrich 98 118 142 0.966 Munster
Dimethyltrisulfide H3 C–S–S–S–CH3 Acros 99 126 165 1.198 Fresh onion
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reach equilibrium. It was verified that the system studied here had reached equilibrium after a 24-h period. 3.3. Membrane type and module Usually, in liquid–liquid extraction applications, the lowest membrane mass transfer resistance is obtained when the membrane pores are filled with the fluid in which the solute is the most soluble. The sulfur compounds studied are mainly hydrophobic, and consequently a hydrophobic microporous membrane was chosen in this study. A cross-flow hollow fiber module of type Liquicel® X-40, made of polypropylene Celgard® hollow fibers, was supplied by Hoechst Celanese. Its main characteristics are given in Table 2. The extractions were operated using the module in counter-current configuration, with the feed aqueous phase and solvent phase being introduced at each opposite extremity of the module. The cross-flow mode is generated by a solid, non-porous wall placed mid-length (baffle) on the shell side of the module. The shell side fluid enters axially at one end of the module, and is forced to cross the fiber bundle from the inside to the outside, in order to go round the wall. In the second half of the module, the fluid goes back to the module axis to exit at its other extremity.
Table 2 Main characteristics of the X-40 Liquicel hollow fiber module Type of fibers Type of potting Porosity (%)a Tortuosityb Pore diameter (m)a Number of fibersa Effective length of fibers (mm)c Fiber inside diameter (m)a Fiber wall thickness (m)a Effective area based on the fiber outside diameter (m2 )a Effective area based on the fiber inside diameter (m2 )c Inside diameter of the hollow fiber bundle (mm)d Outside diameter of the hollow fiber bundle (mm)d Shell side geometric void fractionc
Polypropylene Polyethylene 25 2 0.03 10000 146 200 50 1.4 0.92 12 49 0.40
a
Given by Hoechst Celanese. Estimated by Kiani et al. [18]. c Estimated. d Measured. b
3.4. Pilot-scale extractions A systematic approach was adopted to carry out pilot-scale extractions, with the equipment shown in Fig. 1.
Fig. 1. Schematic view of the pilot-scale extraction device: (1) feed phase reservoir; (2) solvent phase reservoir; (3) hollow fiber module; (4) pumps; (5) by-pass valves and (6) sampling points.
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In all of the experiments, the system was set up as follows. • A synthetic aqueous solution, modeling industrial wastewater, flowed inside the fiber lumen, with volumetric flow rates varying between 100 and 2800 ml min−1 . The solution, at a volume of 5 l, contained the three sulfur compounds at 20 l l−1 each. • The solvent phase was one liter of pure n-hexane flowing on the shell side of the module, at volumetric flow rates in the range of 100–700 ml min−1 . • The minimum pressure difference between the two phases was systematically adjusted between 0.5 and 1 bar in order to prevent the dispersion of one phase into another. It was not monitored accurately since many authors have shown it has no influence on mass transfer [18,25]. This transmembrane pressure is much lower than the critical value characteristic of the membrane used, about 4 bar. • Both phases were totally recycled into their own feed reservoirs. • Both phases were regularly sampled at the bypasses and analyzed by gas chromatography in order to monitor the mass transfers. • All materials making up the extraction unit, except the membrane itself, are made of glass, stainless steel and Teflon® in order to avoid adsorption phenomena. The determination of experimental overall mass transfer coefficients was based on a method proposed by several authors [26,27] and given in Appendix A. This relation was established for counter-current parallel flow modules with total recycling of both aqueous and organic phases. No simple relationship has been written for the real cross-flow module used in this study, because of the complexity of the shellside fluid flow. Thereby, the mass transfer coefficients estimated are apparent mass transfer coefficients, equivalent to pure parallel flow geometry. 3.5. Analytical methods Organic samples were analyzed by gas chromatography using a Carlo Erba GC 5300 chromatograph equipped with a Chrompack CP-SIL capillary column at 50◦ C and a flame ionization detector (FID). Aqueous samples were analyzed by gas chromatography
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as well, using a Chrompack CP 9001 chromatograph equipped with a Hewlett-Packard HP Wax column at 80◦ C and a FID. 3.6. Pervaporation experiments Pervaporation experiments have been performed in a pilot plant [28]. The membrane was a 100 m thick polyetherblockamide (PEBA) membrane supplied by GKSS (Forschugszentrum Geesthacht Gmbh, D-2054 Geesthacht, Germany). Its effective area was 0.1 m2 . The plate and frame module supplied by GFT (Sulzer Chemtech, Membrantechnik D-66540 Neunkirchen, Germany) was modified in order to neglect polarization concentration phenomena in the feed phase [29]. Experiments were carried out at 50◦ C. The feed phase, containing the three sulfur compounds each at 20 ppm (l l−1 ), was flowing at 220 l h−1 , which corresponds to a mean Reynolds number of 600. The total permeate pressure was 350 Pa.
4. Results and discussion 4.1. Partition coefficients The determination of theoretical (Eq. (3)) as well as experimental (Appendix A) overall mass transfer coefficients assumes that the partition coefficients are known. The partition coefficient at equilibrium of each solute between water and n-hexane at 25◦ C was determined for three different concentrations; the range of initial organic concentrations (from 500 to 2000 ppm) was chosen to cover the range of equilibrium aqueous concentrations obtained during extraction experiments. The equilibrium solvent phase concentration as a function of the equilibrium aqueous phase concentration is shown for each solute in Fig. 2. Experimental results show a linear increase in the equilibrium aroma concentration in the organic phase with the equilibrium aqueous concentration. The partition coefficients of the three sulfur compounds between water and n-hexane were constant in the concentration range studied, which can be considered as the infinite dilution range. The values obtained here were then usable in Eq. (3) whatever the exact concentrations at the aqueous–organic inter-
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The observed partition coefficients for these systems were respectively 70 and 0.4. The large values of sulfur compound partition coefficients between water and hexane ensure nearly total extraction from water. They belong to the highest values observed in the literature because of the high hydrophobicity and volatility specific to these aroma compounds. 4.2. Effect of aqueous and organic flow rates on mass transfer kinetics
Fig. 2. Equilibrium curves of sulfur aroma compound partitioning between water and n-hexane at 25◦ C: (䊉, 䊊) DMDS; (䊏, 䊐) MTB and (䉱, ) DMTS. Partition coefficients were determined for each solute alone in solution (closed symbols) and for the mixture of all three (open symbols).
face during the extraction. In order to detect possible interaction between the solutes, an experiment was carried out from a solution containing the three sulfur compounds together in hexane, at a concentration of 1000 ppm each. Results have been plotted (open symbols) in Fig. 2. No significant difference was observed compared to the partition coefficient values obtained with binary solutions. Thus, neither interactive nor competitive effects exist in the mixture studied, and the extraction runs were always performed with the mixture of the three aroma compounds. The partition coefficient values, resulting from the slope obtained by linear regression of the data (Fig. 2), are presented in Table 3 with the regression coefficients. The literature covers very different systems with a wide range of partition coefficients, usually between 10−1 and 103 . For example, D’Elia et al. [26] studied the extraction from water of p-nitrophenol with amyl acetate and of acetic acid with methyl amyl ketone. Table 3 Partition coefficients of sulfur compounds between water and n-hexane at 25◦ C Solute
DMDS
MTB
DMTS
Partition coefficient Correlation coefficient
92.2 0.994
120.6 0.971
564.8 0.960
The flow rates of both phases were alternately modified, with all the other parameters kept constant. Fig. 3 shows the MTB extraction kinetic curves obtained for different aqueous phase flows (a) and different solvent phase flows (b). The tube side (aqueous phase) Reynolds numbers ranged from 1.3 to 32, and the shell side (solvent phase) Reynolds numbers ranged from 0.5 to 3.2. The aqueous flow rate had a significant influence on mass transfer whereas the organic phase flow rate did not seem to have any effect. Much of the mass transfer resistance was due to the aqueous boundary layer thickness. In other terms, the diffusion of solutes from the bulk aqueous phase up to the aqueous–organic interface, through the boundary layer, was the slowest step, controlling the whole extraction operation. It was here necessary to determine the diffusivities of the sulfur compounds in water and in n-hexane. The Wilke–Chang equation [30], recommended in the case of diluted aroma compounds [31], was used. Table 4 summarizes the results. The predominant resistance of the aqueous boundary layer, observed in Fig. 3, was due to the combined effects of two factors. The diffusivities of the three solutes are higher in n-hexane than in water (factor 4), notably due to viscosity differences between water (0.9 mPa s at 25◦ C) and n-hexane (0.3 mPa s), thus increasing the relative importance of the aqueous boundary layer with respect to the organic boundary layer. Moreover, the high partition coefficient values minimize the importance of the two diffusion steps that take place in the solvent phase (see Eq. (3)). Sulfur compounds were extracted efficiently, with recovery rates of over 90% for all three of them, and up to ≥99% for DMTS in 60–90 min of extraction. The corresponding organic concentrations were well below the maximum solubility of the sulfur compounds in
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Table 4 Diffusivities of the sulfur compounds at 25◦ C, based on the Wilke–Chang equation
(×10−6
m3
mol−1 )a
Critical volume Molar volume of solute at its normal boiling temperature (×10−6 m3 mol−1 )b
Diffusion coefficient (×10−10 m2 s−1 ) In n-hexane In water a b
DMDS
MTB
DMTS
255.5 95
375.5 142
309.5 116
44.4 10.9
34.9 8.6
39.4 9.7
Joback modification of Lydersen’s method. Tyn and Calus method.
It should be mentioned that contacting was performed without any leakage of one phase into another. The interface was perfectly stable, even at the beginning of extractions when mass transfers were the highest. 4.3. Prediction of mass transfer coefficients
Fig. 3. MTB extraction kinetics: solvent phase concentration as a function of time (T = 25 ± 1◦ C): (a) different aqueous phase flows (Qorg = 8.5 × 10−6 m3 s−1 ) and (b) different organic phase flows (Qaq = 1.2 × 10−5 m3 s−1 ).
n-hexane. Then, due to high partition coefficients, a single liter of n-hexane may be used to efficiently treat a consequently larger amount of aqueous phase before being saturated.
A resistance-in-series model (Eq. (3)) was used to predict overall mass transfer coefficients by estimating three local mass transfer coefficients. The Lévêque equation used to determine the tube side mass transfer coefficient was chosen because of its convenient shape and also because it has been often used by many other authors. The relation given by Kiani et al. [18] is used widely in the literature for determining mass transfer coefficients of microporous membranes of different types and properties. Eq. (7) proposed by Prasad and Sirkar [21] and Eq. (8) proposed by Wickramasinghe et al. [22] were used to predict shell side mass transfer coefficients from two different approaches. Eq. (7) is consistent with the method used for determining experimental overall mass transfer coefficients (see Appendix A), which is based on the assumption of parallel flows. Eq. (8) may correspond more closely to the actual shell side flow, radially across the fiber bed. Mass transfers predicted by either the parallel flow model (short dash) or the cross-flow model (solid lines) are given in Fig. 4, with regard to experimental data as a function of the aqueous Reynolds number. Diffusivity is the only parameter specific to each solute taken into account when calculating local mass transfers (Eqs. (4)–(6)). As diffusivities were almost equal for the three solutes (Table 4), the
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Fig. 4. Predicted and experimental overall mass transfer coefficients as a function of aqueous flow rates and Reynolds number: (䊉) DMDS; (䊏) MTB; (䉱) DMTS; (—) Eqs. (4), (5) and (8) and (− − −) Eqs. (4), (5) and (7). T = 25 ± 1◦ C, Qorg = 8.5 × 10−6 m3 s−1 .
resistance-in-series model predicted a very slight difference in behavior between them in terms of local mass transfer. The partition coefficients were significantly different for the three solutes, but all three were high enough to have the same effect on mass transfer prediction, i.e. to minimize the terms relative to the organic phase in Eq. (3). The values of experimental overall mass transfer coefficients varied from 5 × 10−6 to 3 × 10−5 m s−1 . These values are of the same order of magnitude as those given in the literature for different systems. Using a microporous hydrophobic flat membrane, Kiani et al. [18] extracted acetic acid from water into
methyl isobutyl ketone (MIBK) with mass transfer coefficients of up to 2 × 10−5 m s−1 . Using hollow fiber modules, Prasad and Sirkar [32] obtained values of 1 × 10−5 and 4 × 10−6 m s−1 for the extraction of 4-methylthiazole and 4-cyanothiazole, respectively, from water with either toluene or benzene. In Fig. 4, the mass transfer coefficients predicted by the Lévêque solution and the correlation of either Prasad and Sirkar (short dash) or Wickramasinghe (solid lines) overestimated experimental data for both sulfides, but corresponded quite well to those of MTB. The relative vertical dispersion of experimental data suggests the existence of an experimental,
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non-reproducible, error. The nature and physicochemical properties of the solutes studied make them difficult to handle; since they are aroma compounds, they are all highly volatile, leading to possible losses during operations such as preparing or loading the feed solutions. Many precautions were taken to minimize experimental errors; the whole extraction unit was made of non-adsorbent materials; aroma compounds solutions were always prepared in hermetically sealed vessels, and handled with great care. Gas chromatography analysis reproducibility was verified. A mass balance in the unit was systematically checked after extraction experiments. Despite these precautions, the particularly low feed phase concentrations (about 0.02 kg m−3 ) affected the overall reproducibility of experiments. Another aspect can be considered to describe the results obtained and explain the differences between experimental data and predicted values of overall mass transfer coefficients. The method used to determine experimental overall mass transfer coefficients (see Appendix A) is valid for a pure parallel counter-current configuration. This equation is based on integrating local mass balances where the two flows on both sides of the membrane wall are strictly parallel. Such an expression is not available for cross-flow modules because the shell side flow is in constant evolution from one end of the module to the other. However, the relative dimensions of the module used in this study (length over radius is about 6) should not confer, on the shell side fluid, overall movement differing excessively from axial flow. Thereby, in a first approach, the equation given in Appendix A was considered to be valid in our case, but a small error was most likely made at this step. The use of the Prasad and Sirkar equation, or any other equation based on axial flow, may also be insufficiently adapted to the cross-flow operating mode. Actually, the choice of a relationship for predicting shell side mass transfer was not very important here, since the resistance-in-series model predicted that the operation would be controlled by diffusion in the aqueous boundary layer. Indeed, whatever the shell side relationship, local aqueous mass transfer resistance represented at least 97% of the whole resistance. However, these two different approaches provided very different Sherwood numbers. The shell side flow characterization of cross-flow
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hollow fiber modules is a complex problem and is being considered by several teams. Efforts must be made in this domain in order to envisage a scale-up operation. Because of the predominance of aqueous boundary layer resistance, the choice of the tube side correlation was particularly important here. The Lévêque solution is a limiting case of the more general Graetz solution, valid for large Reynolds (or Graetz) numbers. When mentioned in the literature, the limit of its applicability is not so clear: for example, Graetz numbers thresholds as different as 4 [20], 25 [33] and 400 [21] can be found in different articles. In any case, the Lévêque solution is known to diverge from the Graetz solution and to overestimate experimental data at low flows. This can be observed in Fig. 5, representing the Lévêque solution (short dash) and the Graetz solution (solid lines) as functions of the tube side modified Graetz number (Gr = Re Sc(dt in /)). The Lévêque solution does not seem to be applicable for Graetz numbers of below 10. Assuming that aqueous resistance controls the whole operation, experimental tube side Sherwood numbers calculation was based on the overall mass transfer coefficient instead of the unavailable aqueous mass transfer coefficient exp (Shtube = (ktube dt in /D) ∼ = (Koverall dt in /D)). Experimental data was spread on both sides of the limit Gr = 10, leading to inaccuracies for the slowest velocities, as suggested by Fig. 5. However, the complexity of
Fig. 5. Predicted and experimental Sherwood number as a function of the tube side Graetz number: (䊉) DMDS; (䊏) MTB; (䉱) DMTS; (—) Graetz solution and (− − −) L´evêque solution.
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the Graetz solution is a clear limit to its use, compared to the Lévêque solution and similar ones. Moreover, since high aqueous Reynolds numbers are desirable for performing fast extractions, the situation for which the Lévêque solution is not applicable corresponds to less favorable conditions. Several authors have pointed out problems of flow characterization, particularly in the case of high density commercial modules [34–36]. Indeed, the high area per volume offered by hollow fiber modules is one of their main advantages. The tendency is thereby to minimize the fiber diameter and maximize their number in the bundle. For this reason, the modules offering the highest theoretical area per volume are also the tightest. The consequence is the occurrence of dead zones, backmixing, bypassing and channeling, especially on the shell side, causing irregular mass transfer along the module. This kind of error, due to the randomness of the number and size of such irregular zones, may not be reproducible and therefore not predictable. However, Seibert et al. [37] have empirically related the fraction of fluid bypassing the hollow fiber bundle to the shell side fluid velocity for a commercial-scale membrane extractor. Moreover, the tube side mass transfer theory is based on the assumption that every tube is perfectly cylindrical. In spite of this, several authors have shown that tube side flow is often not uniform [38]. This drawback has been particularly verified in the case of commercially available modules, but seems to be attenuated when using carefully handmade modules. A major consequence is that, at low flows, the Lévêque solution overestimates experimental mass transfer coefficients [34]. This helps explain the divergence, at low flows, between the resistance-in-series model, that includes the Lévêque solution for tube side mass transfer coefficient prediction, and our experimental data. Moreover, linear regression of log(Kov ) as a function of log(Reaq ) gave gradients in the range of 0.35–0.40. These values, slightly higher than the value of 0.33 given by the Lévêque equation, tend to show the existence of local microturbulences in the tubes, probably due to wall irregularities. Indeed, if real turbulence occurred, these values would be much higher. Finally, the regular use of organic solvents such as n-hexane may alter membrane integrity. Certain fibers may lose their straightness or cylindricity, leading to uneven flows in the bundle.
Table 5 Comparison of mass transfer fluxes obtained experimentally by membrane-based solvent extraction (25◦ C) and by pervaporation (50◦ C), for two different initial driving forces Separation technique
Mass transfer flux (g h−1 m−2 )
Solute
Feed concentration (ppm)
Pervaporation
Membrane-based solvent extraction
DMDS
5 20
0.04 0.56
0.29 1.87
MTB
5 20
0.13 0.78
0.30 1.99
DMTS
5 20
0.26 2.03
0.34 2.60
4.4. Comparison of the performances with other techniques Organophile pervaporation is a particularly well adapted technique for separating volatile molecules from very dilute solutions. The PEBA membrane used is known to be adapted to the extraction of high-boiling compounds [28]. A temperature of 50◦ C was chosen in order to obtain a good balance between fluxes and selectivity. The total pressure on the downstream side of the membrane was equal to 350 Pa, corresponding to the minimum value obtainable with the pilot plant. The performances obtained with membrane-based solvent extraction and pervaporation were compared. Table 5 presents the mass transfer fluxes obtained by these two techniques for two different initial feed phase concentrations, 5 and 20 ppm. Experimental fluxes were of the same order of magnitude, and even slightly higher for membrane-based solvent extraction. The greatest difference was observed for the extraction of DMDS at 5 l l−1 : the flux obtained with membrane-based solvent extraction was more than seven times higher than that obtained with pervaporation. It is important to note that the solvent extraction experiments were carried out at normal temperature, whereas pervaporation experiments were performed at 50◦ C in optimized conditions. Since pervaporation efficiency increases with temperature (due to the exponential increase in flux with temperature), the difference between the two techniques would be even greater if both techniques had been carried out at 25◦ C. Pervaporation is
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very selective because the molecular interactions with the membrane differ from one solute to another. In membrane-based solvent extraction, the only role of the membrane is to physically support and stabilize the interface between the feed phase and the stripping phase; it plays no selective role and thereby mass transfer fluxes are similar for the three solutes. Moreover, the ratio between mass transfer fluxes at two different concentrations is the same in membrane-based solvent extraction, whatever the solute extracted; this characteristic is important because it makes it easier to scale-up the operation. With a feed phase concentration of 20 ppm, membrane-based solvent extraction provided fluxes of 1.87–2.60 g h−1 m−2 for the three aroma compounds, whereas only DMTS was extracted by pervaporation with a similar efficiency. Both the DMDS and MTB fluxes were lower than 0.8 g h−1 m−2 . Membrane-based solvent extraction may be a suitable separation technique for the treatment of very diluted solutions such as those studied here. This is of particular interest for aroma compound recovery as well as for pollution control. The length of a transfer unit (LTU) based on the aqueous phase (Eq. (11)) was calculated from the experimentally determined mass transfer coefficients K, the aqueous phase velocity v and the interfacial area a LTU =
v Ka
(11)
The interfacial area available in the contactor used is about 2400 m2 m−3 . Reed et al. [39] gave the following values for several extraction systems: from 3 to 30 m2 m−3 , free dispersion columns; from 30 to 300 m2 m−3 , packed/trayed columns; from 160 to 500 m2 m−3 , mechanically agitated columns; from 1600 to 6600 m2 m−3 , membrane contactors. It is clear that the very high value of 2400 m2 m−3 is one of the main advantages of such membrane contactors. The lowest LTU obtained in this study were 37, 30 and 50 cm for DMDS, MTB and DMTS, respectively. These values correspond to the lowest velocities applied to the aqueous phase, since v increases faster than Ka in Eq. (11). The effective length of the module, about 15 cm, was shorter than the lowest LTU reached, indicating that the residence times of the solutions in the module were not long enough for the system to reach equilibrium. In other terms, the solvent leaves the module less loaded than its
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capacity, and the feed phase less stripped than it could be. Thus, by using a longer module or several modules in series, the mass transfer rate would be greater for equivalent flow rates. LTU values for the recovery of sulfur aroma compounds are unavailable in the literature, but they can be compared to other systems. Prasad and Sirkar [32] extracted 4-methylthiazole and 4-cyanothiazole from water, using a hydrophobic membrane and either toluene or benzene as the stripping solvent. These authors obtained the remarkably low LTU values of 0.03–0.15 m. The values of LTU obtained in our work were not excessively low, but were still perfectly acceptable. The use of a more porous membrane of the same size, for example, would provide a greater contact area, hence improving the mass transfer coefficients and LTU values.
5. Conclusion Three sulfur aroma compounds were extracted by n-hexane from a dilute aqueous solution representing a food industry wastewater. The operation was carried out without dispersion of one phase into another, using a hydrophobic hollow fiber module. The chosen extraction operating mode consisted in flowing the feed aqueous phase through the tubes, and the solvent phase through the shell, both phases being totally recycled into their own reservoirs during the experiments. The three volatile, hydrophobic sulfur compounds presented remarkably high partition coefficients between water and n-hexane. As a consequence, the aroma compounds were almost totally recovered (extraction yields of 90–99%). The use of n-hexane in the food industry, although authorized, may be a problem in terms of food security. Many other extracting solvents may be used instead of n-hexane, such as natural vegetable oils. According to theory, the controlling step appeared to be the diffusion of solutes from the bulk aqueous phase towards the aqueous–organic interface. Three equations were taken from the literature in order to predict local mass transfer coefficients. A resistance-in-series model was used to estimate overall mass transfer coefficients, with moderate accuracy. The Lévêque solution, chosen to estimate tube side mass transfer coefficients, overestimated experimental observations, particularly for the lowest aqueous
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phase velocities. For Graetz numbers of lower than 10, using the exact Graetz solution may be preferable. Fluxes of about 2–2.5 g h−1 m−2 were obtained experimentally from dilute feed phases (20 l of each solute per liter of solution). A comparison with pervaporation showed that membrane-based solvent extraction provided much higher fluxes in most cases, and was less selective than pervaporation between the three chosen solutes. The operation appeared to be efficient and easily carried out, without any dispersion of one fluid into another. Its modularity is an additional advantage that should contribute to its industrial-scale development. A limit to further development is clearly found in shell side hydrodynamics and mass transfer characterization for hollow fiber modules. The high density of fibers on the shell side, which creates the high area per volume offered by these modules, is also a source of uncontrolled phenomena: bypassing, backmixing, dead zones, channeling. Many teams are working to increase knowledge on and model these phenomena, and we are planning to study this aspect by flowing the aqueous feed phase on the shell side of the module. Lastly, the study will be completed with extraction experiments carried out on real wastewaters, which are highly complex solutions, notably to test the selectivity of the technique with respect to other solutes present in the feed phase as well.
Acknowledgements The authors are indebted to A.M. Wall (the INRA’s translation unit, Jouy en Josas) for revising the English version of the manuscript.
Appendix A. Determination of experimental overall mass transfer coefficients This method was proposed by D’Elia et al. [26]. It comes from integrating local mass balances for a solute between the feed phase and the stripping phase, for a parallel flow counter-current configuration. It is valid for systems in which both the aqueous and organic phases are recycled into their feed reservoirs. This particularity leads to a driving force Ci that decreases with time.
Ciini B + C exp A = t (1 − exp A) ln D Ci
where (Ciiniorg /mi ) − Ciiniaq Ciini = Ci ((Ciiniorg /mi ) − Ci aq ) +(Vaq /mi Vorg )(Ciiniaq − Ci aq ) A = −Ki S(B − C) B=
1 Qaq
C=
1 mi Qorg
D=
1 1 + Vaq mi Vorg
References [1] M. Ramel, Emissions d’odeurs liées au traitement des effluents dans les industries agro-alimentaires, Industries Alimentaires et Agricoles, 116 (10) (1999) 62–70. [2] C.J. Mussinan, M.E. Keelan, Sulfur compounds in foods: an overview, in: C.J. Mussinan, M.E. Keelan (Eds.), Sulfur Compounds in Foods, ACS Symposium Series, 1994, pp. 1–6. [3] A. Baudot, I. Souchon, N. Martin, M. Marin, Application de la pervaporation au traitement d’effluents des industries alimentaires, Industries Alimentaires et Agricoles 115 (10) (1998) 17–26. [4] N. Martin, C. Berger, C. Le Du, H.E. Spinnler, Aroma compound production in cheese curd by coculturing with selected yeast and bacteria, J. Dairy Sci., 2000, in press. [5] H.B. Ding, P.W. Carr, E.L. Cussler, Racemic leucine separation by hollow-fiber extraction, AIChE J. 38 (10) (1992) 1493–1498. [6] G.T. Frank, K.K. Sirkar, Alcohol production by yeast fermentation and membrane extraction, Biotechnol. Bioeng. Symp. 15 (1985) 621–631. [7] G.T. Frank, K.K. Sirkar, An integrated bioreactor-separator: in situ recovery of fermentation products by a novel membrane-based dispersion-free solvent extraction technique, Biotechnol. Bioeng. Symp. 17 (1986) 303–316. [8] R. Basu, K.K. Sirkar, Hollow fiber contained liquid membrane separation of citric acid, AIChE J. 37 (3) (1991) 383–393. [9] L. Dahuron, E.L. Cussler, Protein extractions with hollow fibers, AIChE J. 34 (1) (1988) 130–136. [10] R. Prasad, K.K. Sirkar, Hollow fiber solvent extraction of pharmaceutical products: a case study, J. Membr. Sci. 47 (1989) 235–259. [11] Z.-F. Yang, A.K. Guha, K.K. Sirkar, Novel membrane-based synergistic metal extraction and recovery process, Ind. Eng. Chem. Res. 35 (1996) 1383–1394.
F.X. Pierre et al. / Journal of Membrane Science 187 (2001) 239–253 [12] C.H. Yun, R. Prasad, A.K. Guha, K.K. Sirkar, Hollow fiber solvent extraction removal of toxic heavy metals from aqueous waste streams, Ind. Eng. Chem. Res. 32 (1993) 1186–1195. [13] C.H. Yun, R. Prasad, K.K. Sirkar, Solvent extraction of priority organic pollutants using hollow fiber membranes, in: Proceedings of the AIChE National Meeting, Philadelphia, PA, 21–23 August 1989. [14] C.E. Fabre, P.J. Blanc, A. Marty, G. Goma, I. Souchon, A. Voilley, Extraction of 2-phenylethyl alcohol by techniques such as adsorption, inclusion, supercritical CO2 , liquid–liquid and membrane separations, Perfumer and Flavorist 21 (1996) 27–39. [15] H.E. Spinnler, L. Dufosse, I. Souchon, A. Latrasse, C. Piffaut, A. Voilley P. Delest, Production de ␥-décalactone par bioconversion, Patent FR 53939 C (1993). [16] B.W. Reed, R. Klassen, A.E. Jansen, J.J. Akkerhuis, B.A. Bult, F.I.H.M. Oesterholt, Removal of hydrocarbons from wastewater by membrane extraction, in: Proceedings of the AIChE Spring National Meeting, Atlanta, GA, 17–21 April 1994. [17] R. Prasad, K.K. Sirkar, Membrane-based solvent extraction, in: W.S.W. Ho, K.K. Sirkar (Eds.), Membrane Handbook, Chapman and Hall, New York, 1992, pp. 727–763. [18] A. Kiani, R.R. Bhave, K.K. Sirkar, Solvent extraction with immobilized interfaces in a microporous hydrophobic membrane, J. Membr. Sci. 20 (1984) 125–145. [19] R.E. Beck, J.S. Schultz, Hindered diffusion in microporous membranes with known pore geometry, Science 170 (1970) 1302–1305. [20] A. Gabelman, S.-T. Hwang, Hollow fiber membrane contactors, J. Membr. Sci. 159 (1999) 61–106. [21] R. Prasad, K.K. Sirkar, Dispersion-free solvent extraction with microporous hollow-fiber modules, AIChE J. 34 (2) (1988) 177–188. [22] S.R. Wickramasinghe, M.J. Semmens, E.L. Cussler, Hollow fiber modules made with hollow fiber fabric, J. Membr. Sci. 84 (1993) 1–14. [23] P. Schöner, P. Plucinski, W. Nitsch, U. Daiminger, Mass transfer in the shell side of cross flow hollow fiber modules, Chem. Eng. Sci. 53 (13) (1998) 2319–2326. [24] F.W. Welsh, R.E. Williams, The use of vegetable oils to recover compounds from aqueous solutions, J. Chem. Tech. Biotechnol. 46 (1989) 169–178. [25] R. Prasad, R.R. Bhave, A.K. Kiani, K.K. Sirkar, Further studies on solvent extraction with immobilized interfaces in
[26]
[27]
[28]
[29] [30] [31]
[32] [33]
[34]
[35]
[36] [37]
[38]
[39]
253
a microporous hydrophobic membrane, J. Membr. Sci. 26 (1986) 79–97. N.A. D’Elia, L. Dahuron, E.L. Cussler, Liquid–liquid extractions with microporous hollow fibers, J. Membr. Sci. 29 (1986) 309–319. C.J. Tompkins, A.S. Michaels, S.W. Peretti, Removal of p-nitrophenol from aqueous solution by membrane-supported solvent extraction, J. Membr. Sci. 75 (3) (1992) 277–292. A. Baudot, I. Souchon, M. Marin, Total permeate pressure influence on the selectivity of the pervaporation of aroma compounds, J. Membr. Sci. 158 (1999) 167–185. A. Baudot, M. Marin, Dairy aroma compounds recovery by pervaporation, J. Membr. Sci. 120 (1996) 207–220. R.C. Reid, J.M. Prausnitz, B.E. Poling, The Properties of Gazes and Liquids, McGraw-Hill, New York, 1987. T. Lamer, Extraction de composés d’arôme par pervaporation. Relation entre les propriétés physico-chimiques des substances d’arôme et leur transfert à travers des membranes à base de polydiméthylsiloxane, Ph.D. Thesis, 1993, Université de Bourgogne, Dijon, France. R. Prasad, K.K. Sirkar, Hollow fiber solvent extraction: performances and design, J. Membr. Sci. 50 (1990) 153–175. R.M.C. Viegas, M. Rodr´ıguez, S. Luque, J.R. Alvarez, I.M. Coelhoso, J.P.S.G. Crespo, Mass transfer correlations in membrane extraction: analysis of Wilson-plot methodology, J. Membr. Sci. 145 (1998) 129–142. S.R. Wickramasinghe, M. J Semmens, E.L. Cussler, Mass transfer in various hollow fiber geometries, J. Membr. Sci. 69 (1992) 235–250. M.J. Costello, A.G. Fane, P.A. Hogan, R.W. Schofield, The effect of shell side hydrodynamics on the performance of axial flow hollow fiber modules, J. Membr. Sci. 80 (1993) 1–11. M.-C. Yang, E.L. Cussler, Designing hollow-fiber contactors, AIChE J. 32 (1986) 1910–1915. A.F. Seibert, X. Py, M. Mshewa, J.R. Fair, Hydraulics and mass transfer efficiency of a commercial-scale membrane extractor, Sep. Sci. Technol. 28 (1–3) (1993) 343–359. J.K. Park, H.N. Chang, Flow distribution in the fiber lumen side of a hollow-fiber module, AIChE J. 32 (12) (1985) 1937– 1947. B.W. Reed, M.J. Semmens, E.L. Cussler, Membrane contactors, in: R.D. Noble, S.A. Stern (Eds.), Membrane Separations Technology. Principles and Applications, Elsevier, Amsterdam, 1995, p. 474.
Recovery of sulfur aroma compounds using membrane-based solvent extraction F.X. Pierre, I. Souchon∗ , M. Marin UMR Génie et Microbiologie des Procédés Alimentaires, Institut National de la Recherche Agronomique INRA-INA PG, F-78850 Thiverval Grignon, France Received 3 November 2000; received in revised form 15 January 2001; accepted 17 January 2001
Abstract This work focuses on a non-destructive process for recovering valuable aromatic fractions from the food industry’s odorous wastewaters. Non-dispersive solvent extraction of three sulfur aroma compounds, dimethyldisulfide, dimethyltrisulfide and S-methyl thiobutanoate, was carried out from very diluted aqueous solutions representing real effluent. The mass transfer from water to n-hexane was studied using a cross-flow designed hollow fiber membrane contactor. A preliminary study showed high affinity of solutes for n-hexane, with constant partition coefficients at infinite dilution between water and hexane in a 90–560 range. The influence of tube and shell side hydrodynamics on mass transfer was studied, with the aqueous phase on the tube side, and the organic phase on the shell side. The diffusion of solutes from the bulk aqueous phase to the aqueous–organic interface controlled the separation and contributed, under the conditions tested, to more than 97% of the overall mass transfer resistance. A resistance-in-series model overestimated overall mass transfer coefficients. The main explanation is the inaccuracy of the Lévêque correlation used at low Reynolds numbers. The choice of a correlation for predicting mass transfers in the solvent phase did not affect the estimation, since the corresponding mass transfer resistance was negligible. Mass transfer fluxes obtained experimentally by membrane-based solvent extraction were greater for the three aroma compounds than those obtained by pervaporation. © 2001 Elsevier Science B.V. All rights reserved. Keywords: Membrane contactors; Liquid–liquid extraction; Sulfur aroma compounds; Modeling; Pervaporation
1. Introduction Food manufacturers often generate odorous effluents. Recent environmental constraints have led the food industry to focus on this problem with particular attention. Two kinds of approaches can be taken in order to treat odorous effluents. The first one consists of a process for destroying the odor, such as biofiltration, chemical reactions or incineration. ∗ Corresponding author. Tel.: +33-1-30-81-54-86; fax: +33-1-30-81-55-97. E-mail address: [email protected] (I. Souchon).
Since the molecules involved in odorous pollution are also present in numerous food products as flavoring compounds, it would be interesting to consider the non-destructive treatment of odorous effluent. In this second approach, techniques such as adsorption, absorption or membrane processes can be used. The objectives would in this case be to deodorize the effluent and recover a valuable “natural” aromatic fraction in order to compensate for the cost of the deodorization treatment. Sulfur compounds are the main odorous compounds responsible for olfactory pollution, in particular in food manufactures [1]. These sulfur compounds are
0376-7388/01/$ – see front matter © 2001 Elsevier Science B.V. All rights reserved. PII: S 0 3 7 6 - 7 3 8 8 ( 0 1 ) 0 0 3 5 2 - 0
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Nomenclature a A C d dh D DMDS DMTS e Gz k K LTU m ˙ MTB n ppb ppm P Q Re Sc Sh t v v¯ V
area per volume (m2 m−3 ) membrane area in the module (m2 ) concentration (kg m−3 ) diameter (m) hydraulic diameter (m) diffusion coefficient (m2 s−1 ) dimethyldisulfure dimethyltrisulfure membrane thickness (m) Graetz number local mass transfer coefficient (m s−1 ) overall mass transfer coefficient (m s−1 ) fiber length (m) length of a transfer unit (m) mass transfer flux (kg s−1 ) S-methyl thiobutanoate number of fibers in the bundle volumetric part per billion (1 ppb = 1 l m−3 ) volumetric part per million (1 ppm = 1 l l−1 ) liquid–liquid partition coefficient (kg m−3 /kg m−3 ) volumetric flow rate (m3 s−1 ) Reynolds number Schmidt number Sherwood number time (s) fluid velocity (m s−1 ) mean fluid velocity (m s−1 ) volume (m3 )
Greek letters β constant in Eq. (7) ε membrane porosity ε th geometric void fraction φ module packing density η dynamic viscosity (Pa s) µ kinematic viscosity (m2 s−1 ) ρ fluid density (kg m−3 ) τ membrane tortuosity Subscripts aq aqueous b in relative to the bundle inside diameter
b out i lm mb org shell t in t out tube
relative to the bundle outside diameter relative to solute i logarithmic mean membrane organic relative to the shell side of the module relative to the inside of the tubes relative to the outside of the tubes relative to the tube side of the module
Superscripts ∗ equilibrium value eq equilibrium ini initial
often the result of sulfate or sulfur organic compound reduction with anaerobic bacteria. But these compounds are also often naturally present in food products and can be considered as key aroma compounds, particularly in some cheeses and vegetables [2]. For example, cauliflower blanching water, which is generally spread, contains numerous sulfur compounds responsible for the odor [3]. In this former study, three very odorous sulfur compounds were identified: dimethyldisulfide, dimethyltrisulfide and S-methyl thiobutanoate. These molecules are involved in the aromatic quality of numerous French cheeses [4], and the development of an efficient process to recover them from wastewater processes would be very worthwhile. In the food industry, certain constraints must be taken into account when choosing the process to recover aroma compounds from liquid effluents. On the one hand, it is necessary to respect molecular integrity (no high temperature because of the sensitivity of aroma compounds temperature), and on the other hand to have an efficient process which is easy to install and maintain. The use of membrane contactors can be a technical solution taking all these criteria into account. Such membranes can be used to fix an interface between a liquid and either a gas or another liquid. This latter situation represents an alternative to liquid–liquid extraction referred to as membrane-based solvent extraction or non-dispersive solvent extraction. Non-dispersive extraction is performed by inserting a microporous membrane wall between the feed phase and the stripping phase. Using
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microporous membranes, one of the two fluids to be contacted wets the membrane pores and tends to leak on the other side. To prevent the mixing that would occur, slight overpressure must be applied to the other phase in order to stabilize the interface at the pore mouth. In spite of this, the technique remains a concentration-driven process. With respect to conventional liquid–liquid extraction, many advantages of the membrane contactors can be pointed out. There are no loading, flooding or emulsification problems; except for the pumping of the flowing phases, no agitation or moving part is needed. There is no need for density differences, leading to a greater choice of stripping solvents. A drawback occurs when using a membrane because it creates additional resistance that hinders diffusion from one phase to another, thus slowing the separation down. In most cases, the large surface area per volume offered by hollow fiber modules overcomes this disadvantage. Non-dispersive solvent extraction appears to be a competitive technique in different fields, including • chiral separation [5], • extraction of ethanol and organic acids produced by fermentation [6–8], • protein extraction [9], • pharmaceutical applications [10], • metal ion extraction [11], and • extraction of various pollutants from wastewaters [12,13]. The studies cited cover a wide range of partition coefficients (values from about 10−1 to 103 ), various solvents, membrane shapes and flow configurations. Only a few authors have studied the recovery of aroma compounds by membrane–solvent extraction. Fabre et al. [14] have studied the recovery of 2-phenylethylalcohol (rose-like aroma) produced by fermentation. These authors concluded that non-dispersive solvent extraction was a very promising method with respect to other techniques such as adsorption, pervaporation or simple liquid–liquid extraction. Spinnler et al. [15] have registered a patent for the recovery of gamma-decalactone (peach-like aroma) from fermentation broth by non-dispersive solvent extraction. Reed et al. [16] have presented an example of pilot scale membrane-based solvent extraction: the extraction of numerous organic pollutants was carried
241
out from industrial wastewaters from two plants in Holland. However, industrial applications are scarce, due to the finite life of modules as well as the nature of potting adhesives that may be altered by certain organic solvents. The aim of this work was to apply a technology involving microporous membrane contactors to the recovery of sulfur compounds from very diluted aqueous solutions close to food manufacturer waste streams. In order to improve the separation process, a main objective was to improve the understanding of mass transfer in membrane contactors by identifying the controlling step. Thus, extraction experiments were carried out using a microporous hollow fiber membrane and dilute aqueous solutions representing industrial wastewaters. The influences of tube side and shell side hydrodynamics were studied and discussed, and a resistance-in-series model was used to predict mass transfer coefficients. The experimental performances obtained are discussed and compared to those obtained with another membrane technique, pervaporation.
2. Theoretical approach Membrane-based solvent extraction, like solvent extraction in traditional devices, is a concentration-driven operation. As a general rule, the mass flow rate of solute i (m ˙ i ) from one phase to another is defined by Eq. (1) m ˙ i = Ki A(Ci aq − Ci∗aq )
(1)
where K is the overall mass transfer coefficient based on the aqueous phase concentration, A the surface area of the membrane, Caq the concentration in the ∗ the hypothetical concentraaqueous phase, and Caq tion in an aqueous phase in equilibrium with the actual solvent phase. ∗ and C Caq org are related by the partition coefficient (P), that expresses the equilibrium state as eq
Ci org Ci org Pi = ∗ = eq Ci aq Ci aq
(2)
This equilibrium parameter is dependent upon temperature and concentration. Moreover, P is one of the main criteria for choosing the solvent. Indeed, the
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higher the partition coefficient, the more the equilibrium state is displaced towards the solvent phase, and the more complete the extraction. 2.1. The resistance in series model in hollow fibers Mass transfer is always described as a succession of diffusion steps, from the aqueous boundary layer, through the solvent filled membrane pores, to the solvent phase boundary layer. A resistance-in-series model is used to relate the overall mass transfer coefficient (K) to the local mass transfer coefficients (k). When the aqueous phase flows inside the tube of a hydrophobic hollow fiber membrane, and the organic phase flows outside the tube, the resistance-in-series model is expressed by Eq. (3) 1 1 dt in dt in = + + Ki aq Ki aq Pi ki mb dt lm Pi ki org dt out
(3)
where kaq and korg are the local mass transfer coefficients relative to the aqueous and organic boundary layers, respectively, and kmb the local mass transfer coefficient in the membrane; dt in and dt out are the internal and external fiber diameters, respectively, and dt lm their logarithmic mean value [17]. 2.2. Estimation of local mass transfer coefficients Several assumptions have to be made in order to use the simple expression for the membrane mass transfer coefficient estimation proposed by Kiani et al. [18]. • The membrane is symmetrical and its pores completely wetted by one phase. If the transmembrane pressure P is close to a critical pressure Pcr , partial displacement of the pore filling phase may occur, and the estimation of kmb becomes uncertain; • Diffusion of solutes in the pore-filling fluid must be unhindered. Beck and Schultz [19] suggest that diffusion can be considered to be unhindered if the pore diameter is about two orders of magnitude greater than the solute dimensions. Under these conditions, the mass transfer coefficient inside the membrane can be predicted [18] as ki mb =
Di ε τe
(4)
where Di is the diffusivity of the considered solute in the pore-filling fluid, ε and τ membrane porosity and tortuosity, respectively, and e is membrane thickness, replaced by (dt out −dt in )/2 in the case of hollow fibers. Gabelman and Hwang [20] have reported that the Lévêque solution is widely used in the literature to predict tube side mass transfer. The Lévêque solution (Eq. (5)) is a limiting case of the Graetz solution, applicable to a laminar flow in a tube when the Graetz number is high (Gz > 4), with: 0.33 0.33 0.33 dt in Shtube = 1.62 Retube Sc (5) where Re is the Reynolds number (Re = vd/η), Sc the Schmidt number (Sc = η/D), Sh the Sherwood number (Sc = kd/D) and dt in and the inside fiber diameter and length, respectively. Thus, knowing the Sherwood number, the resulting expression of the tube side mass transfer coefficient (ktube ) is 2 0.33 D v ktube = 1.62 (6) dt in A similar approach, based on the relationship between dimensionless groups, is used for estimating shell side mass transfer coefficients in hollow fiber modules. A summary of these equations, for various experimental systems, is given by Gabelman and Hwang [20]. All of them are in the form of Sh = aReb Sc1/3 , a and b often being fitted from experimental observations. Interestingly, the exponent parameter b over the Reynolds numbers varies from one equation to another, approximately from 1/3 to 1 depending on the geometry of the system studied. Actually, the packing density φ of the membrane module is often included in parameter(s) a and/or b, since it greatly influences the Sherwood number. As an example, Prasad and Sirkar [21] studied liquid–liquid extraction with various solutes and solvents, using a hollow fiber module. For pure counter-currents, they suggested predicting the shell side mass transfer coefficients using 0.33 dh Shshell = β(1 − φ)Re0.6 (7) Sc shell where β is 5.85 or 6.1, respectively for hydrophobic or hydrophilic hollow fibers; in the range of 0 < Re < 500 and 0.04 < φ < 0.4.
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Wickramasinghe et al. [22] studied mass transfers of oxygen from water into nitrogen using three different module shapes (axially wound, hollow fiber fabric and a vane module). These authors proposed an average correlation valid for all three designs: 0.33 Shshell = 0.8 Re0.47 shell Sc
(8)
For cross-flow hollow fiber modules, the velocity of the fluid flowing on the shell side, because of its radial component, is diameter-dependent. The determination of a local velocity implies firstly calculating a hydraulic diameter (dh ) and a geometric void fraction of the fiber bed (ε th ) for the shell side space, as suggested by Schöner et al. [23]. Assuming an ideal cross-flow mode, the mean velocity is obtained by integrating the local velocity from the axis to the wall of the module. Once the hydraulic diameter, geometric void fraction and mean velocity are known, it becomes possible to estimate a mean Reynolds number (Eq. (9)) and a Sherwood number (Eq. (10)), expressed as follows: Reshell =
Shshell =
ρ vd ¯ h 2ρQ (db out + db in )ln(db out /db in ) = µ πµ ndt out (9) kshell dh = kshell D
db2 out
− db2 in
− nd2tout
Dndt out
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responsible for the cauliflower blanching water odor. They are particularly hydrophobic and volatile species, with remarkably low limits of detection by the human nose, of the order of a part per billion (ppb). Their main physical properties are reported in Table 1. n-Hexane was chosen as the extracting solvent for several practical reasons. The sulfur aroma compounds are highly soluble in it, but it is almost insoluble in water. It is authorized in food processing and is widely used in aroma compound manufacturers. Knowing that its boiling point is well below those of sulfur compounds studied, the recovery of these solutes may be based on a simple distillation. Another solvent may be chosen considering many other criteria (direct availability, low price, low toxicity, high safety, etc.). As an example, the use of natural vegetable oils, which are usually good solvents for aroma compounds [24], would generate an aromatized oil possibly usable in a food process. Capillary grade (≥98%) n-hexane was supplied by Carlo Erba, and used as received. 3.2. Partition coefficients between water and n-hexane
(10)
3. Materials and methods 3.1. Chemicals The three solutes chosen for this work are those identified by Baudot et al. [3] as being mainly
A known volume of an organic solution of each solute, at different concentration levels, was carefully poured over a known volume of pure water into a separating funnel. After a 24-h settlement period at 25 ± 1◦ C, both phases were carefully separated and analyzed by gas chromatography. Mixing the two phases should be avoided in order to prevent any loss of hexane in the aqueous phase. It can be replaced by a contacting time long enough for the system to
Table 1 Properties of the sulfur compounds studied
Full name Formula Supplier Purity (%) Molecular weight (g mol−1 ) Boiling point at normal pressure (◦ C) Density at 20◦ C (kg l−1 ) Odor
DMDS
MTB
DMTS
Dimethyldisulfide H3 C–S–S–CH3 Aldrich 99 94 116 1.046 Onion, cauliflower
S-methyl butanoate H3 C–CH2 –CH2 –C(=O)–S–CH3 Aldrich 98 118 142 0.966 Munster
Dimethyltrisulfide H3 C–S–S–S–CH3 Acros 99 126 165 1.198 Fresh onion
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reach equilibrium. It was verified that the system studied here had reached equilibrium after a 24-h period. 3.3. Membrane type and module Usually, in liquid–liquid extraction applications, the lowest membrane mass transfer resistance is obtained when the membrane pores are filled with the fluid in which the solute is the most soluble. The sulfur compounds studied are mainly hydrophobic, and consequently a hydrophobic microporous membrane was chosen in this study. A cross-flow hollow fiber module of type Liquicel® X-40, made of polypropylene Celgard® hollow fibers, was supplied by Hoechst Celanese. Its main characteristics are given in Table 2. The extractions were operated using the module in counter-current configuration, with the feed aqueous phase and solvent phase being introduced at each opposite extremity of the module. The cross-flow mode is generated by a solid, non-porous wall placed mid-length (baffle) on the shell side of the module. The shell side fluid enters axially at one end of the module, and is forced to cross the fiber bundle from the inside to the outside, in order to go round the wall. In the second half of the module, the fluid goes back to the module axis to exit at its other extremity.
Table 2 Main characteristics of the X-40 Liquicel hollow fiber module Type of fibers Type of potting Porosity (%)a Tortuosityb Pore diameter (m)a Number of fibersa Effective length of fibers (mm)c Fiber inside diameter (m)a Fiber wall thickness (m)a Effective area based on the fiber outside diameter (m2 )a Effective area based on the fiber inside diameter (m2 )c Inside diameter of the hollow fiber bundle (mm)d Outside diameter of the hollow fiber bundle (mm)d Shell side geometric void fractionc
Polypropylene Polyethylene 25 2 0.03 10000 146 200 50 1.4 0.92 12 49 0.40
a
Given by Hoechst Celanese. Estimated by Kiani et al. [18]. c Estimated. d Measured. b
3.4. Pilot-scale extractions A systematic approach was adopted to carry out pilot-scale extractions, with the equipment shown in Fig. 1.
Fig. 1. Schematic view of the pilot-scale extraction device: (1) feed phase reservoir; (2) solvent phase reservoir; (3) hollow fiber module; (4) pumps; (5) by-pass valves and (6) sampling points.
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In all of the experiments, the system was set up as follows. • A synthetic aqueous solution, modeling industrial wastewater, flowed inside the fiber lumen, with volumetric flow rates varying between 100 and 2800 ml min−1 . The solution, at a volume of 5 l, contained the three sulfur compounds at 20 l l−1 each. • The solvent phase was one liter of pure n-hexane flowing on the shell side of the module, at volumetric flow rates in the range of 100–700 ml min−1 . • The minimum pressure difference between the two phases was systematically adjusted between 0.5 and 1 bar in order to prevent the dispersion of one phase into another. It was not monitored accurately since many authors have shown it has no influence on mass transfer [18,25]. This transmembrane pressure is much lower than the critical value characteristic of the membrane used, about 4 bar. • Both phases were totally recycled into their own feed reservoirs. • Both phases were regularly sampled at the bypasses and analyzed by gas chromatography in order to monitor the mass transfers. • All materials making up the extraction unit, except the membrane itself, are made of glass, stainless steel and Teflon® in order to avoid adsorption phenomena. The determination of experimental overall mass transfer coefficients was based on a method proposed by several authors [26,27] and given in Appendix A. This relation was established for counter-current parallel flow modules with total recycling of both aqueous and organic phases. No simple relationship has been written for the real cross-flow module used in this study, because of the complexity of the shellside fluid flow. Thereby, the mass transfer coefficients estimated are apparent mass transfer coefficients, equivalent to pure parallel flow geometry. 3.5. Analytical methods Organic samples were analyzed by gas chromatography using a Carlo Erba GC 5300 chromatograph equipped with a Chrompack CP-SIL capillary column at 50◦ C and a flame ionization detector (FID). Aqueous samples were analyzed by gas chromatography
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as well, using a Chrompack CP 9001 chromatograph equipped with a Hewlett-Packard HP Wax column at 80◦ C and a FID. 3.6. Pervaporation experiments Pervaporation experiments have been performed in a pilot plant [28]. The membrane was a 100 m thick polyetherblockamide (PEBA) membrane supplied by GKSS (Forschugszentrum Geesthacht Gmbh, D-2054 Geesthacht, Germany). Its effective area was 0.1 m2 . The plate and frame module supplied by GFT (Sulzer Chemtech, Membrantechnik D-66540 Neunkirchen, Germany) was modified in order to neglect polarization concentration phenomena in the feed phase [29]. Experiments were carried out at 50◦ C. The feed phase, containing the three sulfur compounds each at 20 ppm (l l−1 ), was flowing at 220 l h−1 , which corresponds to a mean Reynolds number of 600. The total permeate pressure was 350 Pa.
4. Results and discussion 4.1. Partition coefficients The determination of theoretical (Eq. (3)) as well as experimental (Appendix A) overall mass transfer coefficients assumes that the partition coefficients are known. The partition coefficient at equilibrium of each solute between water and n-hexane at 25◦ C was determined for three different concentrations; the range of initial organic concentrations (from 500 to 2000 ppm) was chosen to cover the range of equilibrium aqueous concentrations obtained during extraction experiments. The equilibrium solvent phase concentration as a function of the equilibrium aqueous phase concentration is shown for each solute in Fig. 2. Experimental results show a linear increase in the equilibrium aroma concentration in the organic phase with the equilibrium aqueous concentration. The partition coefficients of the three sulfur compounds between water and n-hexane were constant in the concentration range studied, which can be considered as the infinite dilution range. The values obtained here were then usable in Eq. (3) whatever the exact concentrations at the aqueous–organic inter-
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The observed partition coefficients for these systems were respectively 70 and 0.4. The large values of sulfur compound partition coefficients between water and hexane ensure nearly total extraction from water. They belong to the highest values observed in the literature because of the high hydrophobicity and volatility specific to these aroma compounds. 4.2. Effect of aqueous and organic flow rates on mass transfer kinetics
Fig. 2. Equilibrium curves of sulfur aroma compound partitioning between water and n-hexane at 25◦ C: (䊉, 䊊) DMDS; (䊏, 䊐) MTB and (䉱, ) DMTS. Partition coefficients were determined for each solute alone in solution (closed symbols) and for the mixture of all three (open symbols).
face during the extraction. In order to detect possible interaction between the solutes, an experiment was carried out from a solution containing the three sulfur compounds together in hexane, at a concentration of 1000 ppm each. Results have been plotted (open symbols) in Fig. 2. No significant difference was observed compared to the partition coefficient values obtained with binary solutions. Thus, neither interactive nor competitive effects exist in the mixture studied, and the extraction runs were always performed with the mixture of the three aroma compounds. The partition coefficient values, resulting from the slope obtained by linear regression of the data (Fig. 2), are presented in Table 3 with the regression coefficients. The literature covers very different systems with a wide range of partition coefficients, usually between 10−1 and 103 . For example, D’Elia et al. [26] studied the extraction from water of p-nitrophenol with amyl acetate and of acetic acid with methyl amyl ketone. Table 3 Partition coefficients of sulfur compounds between water and n-hexane at 25◦ C Solute
DMDS
MTB
DMTS
Partition coefficient Correlation coefficient
92.2 0.994
120.6 0.971
564.8 0.960
The flow rates of both phases were alternately modified, with all the other parameters kept constant. Fig. 3 shows the MTB extraction kinetic curves obtained for different aqueous phase flows (a) and different solvent phase flows (b). The tube side (aqueous phase) Reynolds numbers ranged from 1.3 to 32, and the shell side (solvent phase) Reynolds numbers ranged from 0.5 to 3.2. The aqueous flow rate had a significant influence on mass transfer whereas the organic phase flow rate did not seem to have any effect. Much of the mass transfer resistance was due to the aqueous boundary layer thickness. In other terms, the diffusion of solutes from the bulk aqueous phase up to the aqueous–organic interface, through the boundary layer, was the slowest step, controlling the whole extraction operation. It was here necessary to determine the diffusivities of the sulfur compounds in water and in n-hexane. The Wilke–Chang equation [30], recommended in the case of diluted aroma compounds [31], was used. Table 4 summarizes the results. The predominant resistance of the aqueous boundary layer, observed in Fig. 3, was due to the combined effects of two factors. The diffusivities of the three solutes are higher in n-hexane than in water (factor 4), notably due to viscosity differences between water (0.9 mPa s at 25◦ C) and n-hexane (0.3 mPa s), thus increasing the relative importance of the aqueous boundary layer with respect to the organic boundary layer. Moreover, the high partition coefficient values minimize the importance of the two diffusion steps that take place in the solvent phase (see Eq. (3)). Sulfur compounds were extracted efficiently, with recovery rates of over 90% for all three of them, and up to ≥99% for DMTS in 60–90 min of extraction. The corresponding organic concentrations were well below the maximum solubility of the sulfur compounds in
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Table 4 Diffusivities of the sulfur compounds at 25◦ C, based on the Wilke–Chang equation
(×10−6
m3
mol−1 )a
Critical volume Molar volume of solute at its normal boiling temperature (×10−6 m3 mol−1 )b
Diffusion coefficient (×10−10 m2 s−1 ) In n-hexane In water a b
DMDS
MTB
DMTS
255.5 95
375.5 142
309.5 116
44.4 10.9
34.9 8.6
39.4 9.7
Joback modification of Lydersen’s method. Tyn and Calus method.
It should be mentioned that contacting was performed without any leakage of one phase into another. The interface was perfectly stable, even at the beginning of extractions when mass transfers were the highest. 4.3. Prediction of mass transfer coefficients
Fig. 3. MTB extraction kinetics: solvent phase concentration as a function of time (T = 25 ± 1◦ C): (a) different aqueous phase flows (Qorg = 8.5 × 10−6 m3 s−1 ) and (b) different organic phase flows (Qaq = 1.2 × 10−5 m3 s−1 ).
n-hexane. Then, due to high partition coefficients, a single liter of n-hexane may be used to efficiently treat a consequently larger amount of aqueous phase before being saturated.
A resistance-in-series model (Eq. (3)) was used to predict overall mass transfer coefficients by estimating three local mass transfer coefficients. The Lévêque equation used to determine the tube side mass transfer coefficient was chosen because of its convenient shape and also because it has been often used by many other authors. The relation given by Kiani et al. [18] is used widely in the literature for determining mass transfer coefficients of microporous membranes of different types and properties. Eq. (7) proposed by Prasad and Sirkar [21] and Eq. (8) proposed by Wickramasinghe et al. [22] were used to predict shell side mass transfer coefficients from two different approaches. Eq. (7) is consistent with the method used for determining experimental overall mass transfer coefficients (see Appendix A), which is based on the assumption of parallel flows. Eq. (8) may correspond more closely to the actual shell side flow, radially across the fiber bed. Mass transfers predicted by either the parallel flow model (short dash) or the cross-flow model (solid lines) are given in Fig. 4, with regard to experimental data as a function of the aqueous Reynolds number. Diffusivity is the only parameter specific to each solute taken into account when calculating local mass transfers (Eqs. (4)–(6)). As diffusivities were almost equal for the three solutes (Table 4), the
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Fig. 4. Predicted and experimental overall mass transfer coefficients as a function of aqueous flow rates and Reynolds number: (䊉) DMDS; (䊏) MTB; (䉱) DMTS; (—) Eqs. (4), (5) and (8) and (− − −) Eqs. (4), (5) and (7). T = 25 ± 1◦ C, Qorg = 8.5 × 10−6 m3 s−1 .
resistance-in-series model predicted a very slight difference in behavior between them in terms of local mass transfer. The partition coefficients were significantly different for the three solutes, but all three were high enough to have the same effect on mass transfer prediction, i.e. to minimize the terms relative to the organic phase in Eq. (3). The values of experimental overall mass transfer coefficients varied from 5 × 10−6 to 3 × 10−5 m s−1 . These values are of the same order of magnitude as those given in the literature for different systems. Using a microporous hydrophobic flat membrane, Kiani et al. [18] extracted acetic acid from water into
methyl isobutyl ketone (MIBK) with mass transfer coefficients of up to 2 × 10−5 m s−1 . Using hollow fiber modules, Prasad and Sirkar [32] obtained values of 1 × 10−5 and 4 × 10−6 m s−1 for the extraction of 4-methylthiazole and 4-cyanothiazole, respectively, from water with either toluene or benzene. In Fig. 4, the mass transfer coefficients predicted by the Lévêque solution and the correlation of either Prasad and Sirkar (short dash) or Wickramasinghe (solid lines) overestimated experimental data for both sulfides, but corresponded quite well to those of MTB. The relative vertical dispersion of experimental data suggests the existence of an experimental,
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non-reproducible, error. The nature and physicochemical properties of the solutes studied make them difficult to handle; since they are aroma compounds, they are all highly volatile, leading to possible losses during operations such as preparing or loading the feed solutions. Many precautions were taken to minimize experimental errors; the whole extraction unit was made of non-adsorbent materials; aroma compounds solutions were always prepared in hermetically sealed vessels, and handled with great care. Gas chromatography analysis reproducibility was verified. A mass balance in the unit was systematically checked after extraction experiments. Despite these precautions, the particularly low feed phase concentrations (about 0.02 kg m−3 ) affected the overall reproducibility of experiments. Another aspect can be considered to describe the results obtained and explain the differences between experimental data and predicted values of overall mass transfer coefficients. The method used to determine experimental overall mass transfer coefficients (see Appendix A) is valid for a pure parallel counter-current configuration. This equation is based on integrating local mass balances where the two flows on both sides of the membrane wall are strictly parallel. Such an expression is not available for cross-flow modules because the shell side flow is in constant evolution from one end of the module to the other. However, the relative dimensions of the module used in this study (length over radius is about 6) should not confer, on the shell side fluid, overall movement differing excessively from axial flow. Thereby, in a first approach, the equation given in Appendix A was considered to be valid in our case, but a small error was most likely made at this step. The use of the Prasad and Sirkar equation, or any other equation based on axial flow, may also be insufficiently adapted to the cross-flow operating mode. Actually, the choice of a relationship for predicting shell side mass transfer was not very important here, since the resistance-in-series model predicted that the operation would be controlled by diffusion in the aqueous boundary layer. Indeed, whatever the shell side relationship, local aqueous mass transfer resistance represented at least 97% of the whole resistance. However, these two different approaches provided very different Sherwood numbers. The shell side flow characterization of cross-flow
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hollow fiber modules is a complex problem and is being considered by several teams. Efforts must be made in this domain in order to envisage a scale-up operation. Because of the predominance of aqueous boundary layer resistance, the choice of the tube side correlation was particularly important here. The Lévêque solution is a limiting case of the more general Graetz solution, valid for large Reynolds (or Graetz) numbers. When mentioned in the literature, the limit of its applicability is not so clear: for example, Graetz numbers thresholds as different as 4 [20], 25 [33] and 400 [21] can be found in different articles. In any case, the Lévêque solution is known to diverge from the Graetz solution and to overestimate experimental data at low flows. This can be observed in Fig. 5, representing the Lévêque solution (short dash) and the Graetz solution (solid lines) as functions of the tube side modified Graetz number (Gr = Re Sc(dt in /)). The Lévêque solution does not seem to be applicable for Graetz numbers of below 10. Assuming that aqueous resistance controls the whole operation, experimental tube side Sherwood numbers calculation was based on the overall mass transfer coefficient instead of the unavailable aqueous mass transfer coefficient exp (Shtube = (ktube dt in /D) ∼ = (Koverall dt in /D)). Experimental data was spread on both sides of the limit Gr = 10, leading to inaccuracies for the slowest velocities, as suggested by Fig. 5. However, the complexity of
Fig. 5. Predicted and experimental Sherwood number as a function of the tube side Graetz number: (䊉) DMDS; (䊏) MTB; (䉱) DMTS; (—) Graetz solution and (− − −) L´evêque solution.
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the Graetz solution is a clear limit to its use, compared to the Lévêque solution and similar ones. Moreover, since high aqueous Reynolds numbers are desirable for performing fast extractions, the situation for which the Lévêque solution is not applicable corresponds to less favorable conditions. Several authors have pointed out problems of flow characterization, particularly in the case of high density commercial modules [34–36]. Indeed, the high area per volume offered by hollow fiber modules is one of their main advantages. The tendency is thereby to minimize the fiber diameter and maximize their number in the bundle. For this reason, the modules offering the highest theoretical area per volume are also the tightest. The consequence is the occurrence of dead zones, backmixing, bypassing and channeling, especially on the shell side, causing irregular mass transfer along the module. This kind of error, due to the randomness of the number and size of such irregular zones, may not be reproducible and therefore not predictable. However, Seibert et al. [37] have empirically related the fraction of fluid bypassing the hollow fiber bundle to the shell side fluid velocity for a commercial-scale membrane extractor. Moreover, the tube side mass transfer theory is based on the assumption that every tube is perfectly cylindrical. In spite of this, several authors have shown that tube side flow is often not uniform [38]. This drawback has been particularly verified in the case of commercially available modules, but seems to be attenuated when using carefully handmade modules. A major consequence is that, at low flows, the Lévêque solution overestimates experimental mass transfer coefficients [34]. This helps explain the divergence, at low flows, between the resistance-in-series model, that includes the Lévêque solution for tube side mass transfer coefficient prediction, and our experimental data. Moreover, linear regression of log(Kov ) as a function of log(Reaq ) gave gradients in the range of 0.35–0.40. These values, slightly higher than the value of 0.33 given by the Lévêque equation, tend to show the existence of local microturbulences in the tubes, probably due to wall irregularities. Indeed, if real turbulence occurred, these values would be much higher. Finally, the regular use of organic solvents such as n-hexane may alter membrane integrity. Certain fibers may lose their straightness or cylindricity, leading to uneven flows in the bundle.
Table 5 Comparison of mass transfer fluxes obtained experimentally by membrane-based solvent extraction (25◦ C) and by pervaporation (50◦ C), for two different initial driving forces Separation technique
Mass transfer flux (g h−1 m−2 )
Solute
Feed concentration (ppm)
Pervaporation
Membrane-based solvent extraction
DMDS
5 20
0.04 0.56
0.29 1.87
MTB
5 20
0.13 0.78
0.30 1.99
DMTS
5 20
0.26 2.03
0.34 2.60
4.4. Comparison of the performances with other techniques Organophile pervaporation is a particularly well adapted technique for separating volatile molecules from very dilute solutions. The PEBA membrane used is known to be adapted to the extraction of high-boiling compounds [28]. A temperature of 50◦ C was chosen in order to obtain a good balance between fluxes and selectivity. The total pressure on the downstream side of the membrane was equal to 350 Pa, corresponding to the minimum value obtainable with the pilot plant. The performances obtained with membrane-based solvent extraction and pervaporation were compared. Table 5 presents the mass transfer fluxes obtained by these two techniques for two different initial feed phase concentrations, 5 and 20 ppm. Experimental fluxes were of the same order of magnitude, and even slightly higher for membrane-based solvent extraction. The greatest difference was observed for the extraction of DMDS at 5 l l−1 : the flux obtained with membrane-based solvent extraction was more than seven times higher than that obtained with pervaporation. It is important to note that the solvent extraction experiments were carried out at normal temperature, whereas pervaporation experiments were performed at 50◦ C in optimized conditions. Since pervaporation efficiency increases with temperature (due to the exponential increase in flux with temperature), the difference between the two techniques would be even greater if both techniques had been carried out at 25◦ C. Pervaporation is
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very selective because the molecular interactions with the membrane differ from one solute to another. In membrane-based solvent extraction, the only role of the membrane is to physically support and stabilize the interface between the feed phase and the stripping phase; it plays no selective role and thereby mass transfer fluxes are similar for the three solutes. Moreover, the ratio between mass transfer fluxes at two different concentrations is the same in membrane-based solvent extraction, whatever the solute extracted; this characteristic is important because it makes it easier to scale-up the operation. With a feed phase concentration of 20 ppm, membrane-based solvent extraction provided fluxes of 1.87–2.60 g h−1 m−2 for the three aroma compounds, whereas only DMTS was extracted by pervaporation with a similar efficiency. Both the DMDS and MTB fluxes were lower than 0.8 g h−1 m−2 . Membrane-based solvent extraction may be a suitable separation technique for the treatment of very diluted solutions such as those studied here. This is of particular interest for aroma compound recovery as well as for pollution control. The length of a transfer unit (LTU) based on the aqueous phase (Eq. (11)) was calculated from the experimentally determined mass transfer coefficients K, the aqueous phase velocity v and the interfacial area a LTU =
v Ka
(11)
The interfacial area available in the contactor used is about 2400 m2 m−3 . Reed et al. [39] gave the following values for several extraction systems: from 3 to 30 m2 m−3 , free dispersion columns; from 30 to 300 m2 m−3 , packed/trayed columns; from 160 to 500 m2 m−3 , mechanically agitated columns; from 1600 to 6600 m2 m−3 , membrane contactors. It is clear that the very high value of 2400 m2 m−3 is one of the main advantages of such membrane contactors. The lowest LTU obtained in this study were 37, 30 and 50 cm for DMDS, MTB and DMTS, respectively. These values correspond to the lowest velocities applied to the aqueous phase, since v increases faster than Ka in Eq. (11). The effective length of the module, about 15 cm, was shorter than the lowest LTU reached, indicating that the residence times of the solutions in the module were not long enough for the system to reach equilibrium. In other terms, the solvent leaves the module less loaded than its
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capacity, and the feed phase less stripped than it could be. Thus, by using a longer module or several modules in series, the mass transfer rate would be greater for equivalent flow rates. LTU values for the recovery of sulfur aroma compounds are unavailable in the literature, but they can be compared to other systems. Prasad and Sirkar [32] extracted 4-methylthiazole and 4-cyanothiazole from water, using a hydrophobic membrane and either toluene or benzene as the stripping solvent. These authors obtained the remarkably low LTU values of 0.03–0.15 m. The values of LTU obtained in our work were not excessively low, but were still perfectly acceptable. The use of a more porous membrane of the same size, for example, would provide a greater contact area, hence improving the mass transfer coefficients and LTU values.
5. Conclusion Three sulfur aroma compounds were extracted by n-hexane from a dilute aqueous solution representing a food industry wastewater. The operation was carried out without dispersion of one phase into another, using a hydrophobic hollow fiber module. The chosen extraction operating mode consisted in flowing the feed aqueous phase through the tubes, and the solvent phase through the shell, both phases being totally recycled into their own reservoirs during the experiments. The three volatile, hydrophobic sulfur compounds presented remarkably high partition coefficients between water and n-hexane. As a consequence, the aroma compounds were almost totally recovered (extraction yields of 90–99%). The use of n-hexane in the food industry, although authorized, may be a problem in terms of food security. Many other extracting solvents may be used instead of n-hexane, such as natural vegetable oils. According to theory, the controlling step appeared to be the diffusion of solutes from the bulk aqueous phase towards the aqueous–organic interface. Three equations were taken from the literature in order to predict local mass transfer coefficients. A resistance-in-series model was used to estimate overall mass transfer coefficients, with moderate accuracy. The Lévêque solution, chosen to estimate tube side mass transfer coefficients, overestimated experimental observations, particularly for the lowest aqueous
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phase velocities. For Graetz numbers of lower than 10, using the exact Graetz solution may be preferable. Fluxes of about 2–2.5 g h−1 m−2 were obtained experimentally from dilute feed phases (20 l of each solute per liter of solution). A comparison with pervaporation showed that membrane-based solvent extraction provided much higher fluxes in most cases, and was less selective than pervaporation between the three chosen solutes. The operation appeared to be efficient and easily carried out, without any dispersion of one fluid into another. Its modularity is an additional advantage that should contribute to its industrial-scale development. A limit to further development is clearly found in shell side hydrodynamics and mass transfer characterization for hollow fiber modules. The high density of fibers on the shell side, which creates the high area per volume offered by these modules, is also a source of uncontrolled phenomena: bypassing, backmixing, dead zones, channeling. Many teams are working to increase knowledge on and model these phenomena, and we are planning to study this aspect by flowing the aqueous feed phase on the shell side of the module. Lastly, the study will be completed with extraction experiments carried out on real wastewaters, which are highly complex solutions, notably to test the selectivity of the technique with respect to other solutes present in the feed phase as well.
Acknowledgements The authors are indebted to A.M. Wall (the INRA’s translation unit, Jouy en Josas) for revising the English version of the manuscript.
Appendix A. Determination of experimental overall mass transfer coefficients This method was proposed by D’Elia et al. [26]. It comes from integrating local mass balances for a solute between the feed phase and the stripping phase, for a parallel flow counter-current configuration. It is valid for systems in which both the aqueous and organic phases are recycled into their feed reservoirs. This particularity leads to a driving force Ci that decreases with time.
Ciini B + C exp A = t (1 − exp A) ln D Ci
where (Ciiniorg /mi ) − Ciiniaq Ciini = Ci ((Ciiniorg /mi ) − Ci aq ) +(Vaq /mi Vorg )(Ciiniaq − Ci aq ) A = −Ki S(B − C) B=
1 Qaq
C=
1 mi Qorg
D=
1 1 + Vaq mi Vorg
References [1] M. Ramel, Emissions d’odeurs liées au traitement des effluents dans les industries agro-alimentaires, Industries Alimentaires et Agricoles, 116 (10) (1999) 62–70. [2] C.J. Mussinan, M.E. Keelan, Sulfur compounds in foods: an overview, in: C.J. Mussinan, M.E. Keelan (Eds.), Sulfur Compounds in Foods, ACS Symposium Series, 1994, pp. 1–6. [3] A. Baudot, I. Souchon, N. Martin, M. Marin, Application de la pervaporation au traitement d’effluents des industries alimentaires, Industries Alimentaires et Agricoles 115 (10) (1998) 17–26. [4] N. Martin, C. Berger, C. Le Du, H.E. Spinnler, Aroma compound production in cheese curd by coculturing with selected yeast and bacteria, J. Dairy Sci., 2000, in press. [5] H.B. Ding, P.W. Carr, E.L. Cussler, Racemic leucine separation by hollow-fiber extraction, AIChE J. 38 (10) (1992) 1493–1498. [6] G.T. Frank, K.K. Sirkar, Alcohol production by yeast fermentation and membrane extraction, Biotechnol. Bioeng. Symp. 15 (1985) 621–631. [7] G.T. Frank, K.K. Sirkar, An integrated bioreactor-separator: in situ recovery of fermentation products by a novel membrane-based dispersion-free solvent extraction technique, Biotechnol. Bioeng. Symp. 17 (1986) 303–316. [8] R. Basu, K.K. Sirkar, Hollow fiber contained liquid membrane separation of citric acid, AIChE J. 37 (3) (1991) 383–393. [9] L. Dahuron, E.L. Cussler, Protein extractions with hollow fibers, AIChE J. 34 (1) (1988) 130–136. [10] R. Prasad, K.K. Sirkar, Hollow fiber solvent extraction of pharmaceutical products: a case study, J. Membr. Sci. 47 (1989) 235–259. [11] Z.-F. Yang, A.K. Guha, K.K. Sirkar, Novel membrane-based synergistic metal extraction and recovery process, Ind. Eng. Chem. Res. 35 (1996) 1383–1394.
F.X. Pierre et al. / Journal of Membrane Science 187 (2001) 239–253 [12] C.H. Yun, R. Prasad, A.K. Guha, K.K. Sirkar, Hollow fiber solvent extraction removal of toxic heavy metals from aqueous waste streams, Ind. Eng. Chem. Res. 32 (1993) 1186–1195. [13] C.H. Yun, R. Prasad, K.K. Sirkar, Solvent extraction of priority organic pollutants using hollow fiber membranes, in: Proceedings of the AIChE National Meeting, Philadelphia, PA, 21–23 August 1989. [14] C.E. Fabre, P.J. Blanc, A. Marty, G. Goma, I. Souchon, A. Voilley, Extraction of 2-phenylethyl alcohol by techniques such as adsorption, inclusion, supercritical CO2 , liquid–liquid and membrane separations, Perfumer and Flavorist 21 (1996) 27–39. [15] H.E. Spinnler, L. Dufosse, I. Souchon, A. Latrasse, C. Piffaut, A. Voilley P. Delest, Production de ␥-décalactone par bioconversion, Patent FR 53939 C (1993). [16] B.W. Reed, R. Klassen, A.E. Jansen, J.J. Akkerhuis, B.A. Bult, F.I.H.M. Oesterholt, Removal of hydrocarbons from wastewater by membrane extraction, in: Proceedings of the AIChE Spring National Meeting, Atlanta, GA, 17–21 April 1994. [17] R. Prasad, K.K. Sirkar, Membrane-based solvent extraction, in: W.S.W. Ho, K.K. Sirkar (Eds.), Membrane Handbook, Chapman and Hall, New York, 1992, pp. 727–763. [18] A. Kiani, R.R. Bhave, K.K. Sirkar, Solvent extraction with immobilized interfaces in a microporous hydrophobic membrane, J. Membr. Sci. 20 (1984) 125–145. [19] R.E. Beck, J.S. Schultz, Hindered diffusion in microporous membranes with known pore geometry, Science 170 (1970) 1302–1305. [20] A. Gabelman, S.-T. Hwang, Hollow fiber membrane contactors, J. Membr. Sci. 159 (1999) 61–106. [21] R. Prasad, K.K. Sirkar, Dispersion-free solvent extraction with microporous hollow-fiber modules, AIChE J. 34 (2) (1988) 177–188. [22] S.R. Wickramasinghe, M.J. Semmens, E.L. Cussler, Hollow fiber modules made with hollow fiber fabric, J. Membr. Sci. 84 (1993) 1–14. [23] P. Schöner, P. Plucinski, W. Nitsch, U. Daiminger, Mass transfer in the shell side of cross flow hollow fiber modules, Chem. Eng. Sci. 53 (13) (1998) 2319–2326. [24] F.W. Welsh, R.E. Williams, The use of vegetable oils to recover compounds from aqueous solutions, J. Chem. Tech. Biotechnol. 46 (1989) 169–178. [25] R. Prasad, R.R. Bhave, A.K. Kiani, K.K. Sirkar, Further studies on solvent extraction with immobilized interfaces in
[26]
[27]
[28]
[29] [30] [31]
[32] [33]
[34]
[35]
[36] [37]
[38]
[39]
253
a microporous hydrophobic membrane, J. Membr. Sci. 26 (1986) 79–97. N.A. D’Elia, L. Dahuron, E.L. Cussler, Liquid–liquid extractions with microporous hollow fibers, J. Membr. Sci. 29 (1986) 309–319. C.J. Tompkins, A.S. Michaels, S.W. Peretti, Removal of p-nitrophenol from aqueous solution by membrane-supported solvent extraction, J. Membr. Sci. 75 (3) (1992) 277–292. A. Baudot, I. Souchon, M. Marin, Total permeate pressure influence on the selectivity of the pervaporation of aroma compounds, J. Membr. Sci. 158 (1999) 167–185. A. Baudot, M. Marin, Dairy aroma compounds recovery by pervaporation, J. Membr. Sci. 120 (1996) 207–220. R.C. Reid, J.M. Prausnitz, B.E. Poling, The Properties of Gazes and Liquids, McGraw-Hill, New York, 1987. T. Lamer, Extraction de composés d’arôme par pervaporation. Relation entre les propriétés physico-chimiques des substances d’arôme et leur transfert à travers des membranes à base de polydiméthylsiloxane, Ph.D. Thesis, 1993, Université de Bourgogne, Dijon, France. R. Prasad, K.K. Sirkar, Hollow fiber solvent extraction: performances and design, J. Membr. Sci. 50 (1990) 153–175. R.M.C. Viegas, M. Rodr´ıguez, S. Luque, J.R. Alvarez, I.M. Coelhoso, J.P.S.G. Crespo, Mass transfer correlations in membrane extraction: analysis of Wilson-plot methodology, J. Membr. Sci. 145 (1998) 129–142. S.R. Wickramasinghe, M. J Semmens, E.L. Cussler, Mass transfer in various hollow fiber geometries, J. Membr. Sci. 69 (1992) 235–250. M.J. Costello, A.G. Fane, P.A. Hogan, R.W. Schofield, The effect of shell side hydrodynamics on the performance of axial flow hollow fiber modules, J. Membr. Sci. 80 (1993) 1–11. M.-C. Yang, E.L. Cussler, Designing hollow-fiber contactors, AIChE J. 32 (1986) 1910–1915. A.F. Seibert, X. Py, M. Mshewa, J.R. Fair, Hydraulics and mass transfer efficiency of a commercial-scale membrane extractor, Sep. Sci. Technol. 28 (1–3) (1993) 343–359. J.K. Park, H.N. Chang, Flow distribution in the fiber lumen side of a hollow-fiber module, AIChE J. 32 (12) (1985) 1937– 1947. B.W. Reed, M.J. Semmens, E.L. Cussler, Membrane contactors, in: R.D. Noble, S.A. Stern (Eds.), Membrane Separations Technology. Principles and Applications, Elsevier, Amsterdam, 1995, p. 474.