Journal of Membrane Science 213 (2003) 181–193
Recovery of phenol from aqueous solutions using hollow fibre contactors M.J. González-Muñoz, S. Luque, J.R. Álvarez, J. Coca∗ Department of Chemical and Environmental Engineering, C/Julian Claveria s/n, University of Oviedo, 33071 Oviedo, Spain Received 3 April 2002; received in revised form 6 November 2002; accepted 6 November 2002
Abstract The purpose of this study is to characterise recovery of phenol from an aqueous solution using a hydrophobic polypropylene membrane contactor. The effects of temperature and hydrodynamics on the overall mass transfer coefficient were determined. Integration of the extraction and stripping stages was also carried out thereby allowing removal of more than 99% of the original phenol, while the organic phase is simultaneously regenerated. © 2002 Elsevier Science B.V. All rights reserved. Keywords: Phenol; 1-Decanol; Membrane contactor; Mass transfer coefficient; Extraction-stripping
1. Introduction Phenols are often present in wastewaters from many industrial processes, such as refineries (6–500 mg/l), coking operations (28–3900 mg/l), coal processing (9–6800 mg/l), and manufacture of petrochemicals (2.8–1220 mg/l). Phenols are also the main organic constituents present in condensate streams in coal gasification and liquefaction processes [1,2]. Other sources of waste stream water containing phenols are pharmaceutical, plastics, wood products, paint, and pulp and paper industries (0.1–1600 mg/l) [3–8]. The maximum allowed concentration of phenol in non-chlorinated water is 0.1 mg/l while that in chlorinated water is 0.001–0.002 mg/l. Some standards adopted for water supplies are as low as 1–2 ppb [9–11].
∗ Corresponding author. Tel.: +34-98-510-3443; fax: +34-98-510-3434/3443. E-mail address:
[email protected] (J. Coca).
Phenols are very difficult to degrade, because they are toxic for micro-organisms. Furthermore, phenol reacts with the chlorine used in the treatment of potable water, yielding compounds even more toxic and more resistant to biodegradation than phenol [12]. When the phenol concentration is low (97% (w/w)), trioctylmethylammonium chloride (TOMAC) and trioctylamine (TOA, 95% (w/w)) were supplied by Fluka. All chemicals were used without further purification. 2.2. Experimental apparatus A schematic diagram of the experimental apparatus is shown in Fig. 1. The equipment consists of a polypropylene hollow fibre membrane contactor and
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Fig. 1. Schematic diagram of extraction-stripping apparatus.
two jacketed vessels containing either the aqueous or the organic phases. The hollow fibre membrane module is a Liqui-Cel Extra-Flow phase contactor that contains 10176 hydrophobic polypropylene Celgard® fibres. Each fibre is 15 cm length and has a nominal internal diameter of 240 m, a nominal thickness of 30 m and a pore size of 0.03 m. The effective surface area is 1.4 m2 . The membrane module is provided with a central tube (0.0222 m diameter) and a central baffle to facilitate good distribution of the flow through
the shell side (see Fig. 2). The internal diameter of the casing is 0.0555 m. The aqueous phase is pumped through the lumen of the fibres, and the organic phase flows through the shell side of the module. Transfer of solute from the aqueous feed stream to the organic phase occurs through the pores of the membrane. Since the membrane is hydrophobic, a slight overpressure of the aqueous phase (0.6–0.8 bar) is necessary to stabilise the interface in the pores so as to avoid bulk mixing of the two phases.
Fig. 2. Liqui-Cel Extra-Flow membrane contactor with flow of aqueous phase through the lumens of the hollow fibres.
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Ethylene glycol from a constant temperature bath is pumped through the jacket of each vessel in order to ensure a constant temperature throughout the experiment. In order to carry out the extraction and stripping processes simultaneously, a second membrane module is used. In the process of interest, the organic phase is pumped through the first module. and then, now loaded with solute, is transferred to the shell-side of stripping module. In this second module, the organic phase is contacted with the aqueous stripping solution of NaOH, thereby achieving continuous regeneration of the organic phase. 2.3. Methods 2.3.1. Selection of the extraction system Batch extraction experiments were carried out using several solvents (1-decanol, hexane, toluene, methylcyclohexane, n-heptane, MIBK, Petrosol 15/20 and Petrosol D15/20). In order to study physical extraction by the aforementioned solvents, a known volume (10–20 ml) of aqueous phase, containing 3000 mg/l of phenol, was mixed with the same volume of an organic phase in a screw-capped flask. The flasks were immersed in a constant temperature bath (20 ◦ C) and stirred for at least 4 h, long enough to approach equilibrium. After phase separation, phenol concentrations were measured for both phases. For the solvents providing the highest degrees of extraction, distribution coefficients were measured using different ratios of the volumes of the aqueous and organic phases (with a constant initial concentration of phenol in the aqueous phase). After the physical extraction experiments, the reactive extractant was added to the solvent (diluent) with the highest distribution coefficient at a concentration of 5% (v/v). The aforementioned commercial extractants were tested in the extraction experiments. Sulphuric acid salts of TOA were obtained by mixing 80 l of H2 SO4 (96% (w/w)) with 10 ml of TOA [22]. All experiments in the reactive extractant phase of the work were carried out at 20 ◦ C with a volume ratio of 2 (aqueous/organic). 2.3.2. Membrane contactor experiments The experiments were designed to study the effects of hydrodynamics (Reynolds number and flow pattern,
i.e. co-current or counter-current operation) and temperature (in the range 20–40 ◦ C) on the overall mass transfer coefficient. The effect of the Reynolds number was studied by establishing a specific flow rate for one phase while varying the flow rate of the other phase. The Reynolds number for the shell side was calculated using the expression provided by the manufacturer, which takes into account the geometry of the module: DH vρ Re = (1) µ where DH is the hydraulic diameter, v the average linear velocity and ρ and µ are the density and viscosity of the organic phase, respectively. DH was calculated as follows: 2 (π/4)Dshell D2 DH = 4 (2) = shell Nf π d f N f df where Dshell is the inside diameter of the shell, Nf the number of fibres, and df is the outside diameter of a single fibre. The value of DH in this case is 10−3 m. The average linear velocity is given by Qo (3) v= A where Qo is the volumetric flow rate, and A is the area of the shell through which the organic phase flows, calculated as 2 A = 41 π(Dshell − Dt2 − Nf df2 )
(4)
where Dt is the outside diameter of the central (collection and distribution) tube (see Fig. 2). For the stripping experiments, the effect of the mole ratio of NaOH to phenol on the recovery of phenol was studied to determine the optimum value of this ratio, i.e. the ratio that gives the highest recovery of phenol at the lowest concentration of NaOH. The effect of hydrodynamics during the stripping process has not been studied since the same behaviour as in the aqueous phase during the extraction process is expected. Both the feed and stripping solutions are diluted aqueous streams. Therefore, influencing parameters such us density, viscosity and interfacial tension would not change significantly, and the optimum operating conditions should be similar in both cases. During the membrane contactor experiments, samples of each phase were taken at specific times in order to monitor the time course of the phenol concentration profiles in each phase.
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Once the optimum conditions (hydrodynamics and flow pattern) for the extraction process were established, experiments involving the integrated process were carried out. The combination of the extraction and stripping processes in the two membrane contactors permits one to transfer the solute from an aqueous phase (feed) to a second aqueous phase (stripping) while continuously regenerating the organic phase.
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where A m Ka Qa (1 + V ) 1 − exp − (1 + Q) (7) c= Va (1 + Q) Qa and V =
Va Vo D
(8)
Q=
Qa Qo D
(9)
2.4. Analytical methods
For counter-current flow [26]: Phenol solutions were analysed by HPLC, using a Hewlett Packard Hypersil MOS column. The UV absorbances at 254 and 280 nm were measured using a diode array detector. The mobile phase was a mixture of methanol and water (55/45 (v/v)) for the analysis of phenol in aqueous phases, and methanol/acetonitrile/water (40/40/20 (v/v/v)) for determination of phenol in the organic phase [22]. The flow rate was 1 ml/min in all cases. The NaOH concentration in the stripping phase was determined by titration with potassium hydrogen phthalate using phenolphthalein as the indicator. 2.5. Mass transfer coefficient calculation
ca =
Vc0a ca0 + exp(−c t) 1+V 1+V
where Qa 1 − exp[φ(Q − 1)] (1 + V ) c = Va 1 − Q exp[φ(Q − 1)]
φ=
A m Ka Qa
Jsol = Ka Am (ca − ca∗ )
3.1. Selection of the extraction system
ca =
Vc0a ca0 + exp(−ct) 1+V 1+V
(6)
(12)
The initial phenol concentration in the aqueous phase is ca0 , Va and Vo are the volumes of the aqueous and organic phases, respectively. Qa and Qo are the flow rates of the aqueous feed and organic phase, respectively.
3. Results and discussion
where Jsol is the solute flux, Am the membrane area, ca the concentration of phenol in the aqueous solution at time t, and ca∗ is the concentration of phenol in the aqueous phase in equilibrium with the organic phase at this same time. Through a mass balance, the flux of solute in Eq. (5) can be related to the depletion of solute in the aqueous phase. If one assumes a high degree of mixing in the feed vessels (the solute concentration in the vessel being equal to that entering the module) and a constant distribution coefficient, D, the solute concentration in the aqueous phase may be obtained [16,26,27]. The equation for co-current flow is (a detailed derivation can be found in Appendix A):
(11)
and
The transfer of phenol from the aqueous phase to the organic phase can be described in terms of the overall mass transfer coefficient, Ka : (5)
(10)
For the solvents used in the preliminary extraction experiments the equilibrium concentration of phenol in the organic and aqueous phases are shown in Fig. 3. The solvents that produced the greatest degree of extraction (1-decanol, toluene and MIBK) were further studied to obtain the equilibrium isotherms (20 ◦ C) shown in Fig. 4. In these experiments mass balances closed, in general, within 3%. Two of the most important characteristics of a solvent to be used to extract a particular solute, namely, a high distribution coefficient and a low solubility in water [6,23] are presented in Table 1 for 1-decanol, toluene and MIBK. On the basis of these data, 1-decanol was selected for further study. This solvent has both a high distribution coefficient and very low solubility in water.
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Fig. 3. Distribution of phenol between water (䊐) and several solvents (䊏). Initial phenol concentration in aqueous phase: 3000 ppm. Volume ratio (Va :Vo ) = 1:1.
Solutions of reactive extractants at a concentrations of 5% (v/v) in 1-decanol were then tested to assess their potential for use as extraction solvents for phenol (see Fig. 5). The degree of extraction obtained using a reactive extractant was slightly higher than that ob-
tained using just the diluent (1-decanol) itself. Consequently, use of a reactive extraction system cannot be justified. Hence, all membrane contactor experiments were carried out using pure 1-decanol as the organic solvent (extractant).
Fig. 4. Extraction isotherms at 20 ◦ C: (䊐) methylisobutylketone; (䊉) 1-decanol; () toluene.
M.J. Gonz´alez-Muñoz et al. / Journal of Membrane Science 213 (2003) 181–193 Table 1 Distribution coefficients and solubility in water of some solvents [24,25] Solvent
Distribution coefficient at 20 ◦ C
MIBK 1-Decanol Toluene
83.2 25.4 1.7
Solubility in water (mg/l) at 25 ◦ C 17,000 37 526
3.2. Calculation of mass transfer coefficient The time dependence of the concentration of phenol in the aqueous phase can be expressed (see Eqs. (6)–(12)) as the sum of two terms ca = a + b exp(−ct)
(13)
where a and b depend only on known parameters. The value of c can be obtained from a plot of ln(ca − a) against time, c being the slope of the straight line thus obtained. Then, Ka can be determined from either Eq. (7) or Eq. (11), for co-current or counter-current flow, respectively. For a co-current flow pattern: −Qa Va 1+Q Ka = ln 1 − c (14) (1 + Q)Am Qa 1+V
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3.3. Influence of hydrodynamics and flow pattern The data for the extraction experiments provide very regular and smooth solute concentration profiles (Fig. 6). The overall mass transfer coefficient, Ka , is depicted in Fig. 7a as a function of the Reynolds number of the organic phase (shell side), Reo , for the extraction experiments performed at 20 ◦ C with co-current flow of aqueous and organic phases. The Reynolds number of the organic phase (Reo ) was varied in the range 0.1–0.65 (Qo = 8–50 l/h) while keeping the tube side Reynolds number constant (Rea = 4; Qa = 28 l/h). For these experiments, Ka increased with Reo until an asymptotic value is reached at about Reo = 0.4. When the Reynolds number of the aqueous phase was varied while keeping the shell side Reynolds number constant (Reo = 0.32), no further increase in the overall mass transfer coefficient was observed. Therefore, the system had reached the conditions in which the controlling resistance to mass transfer is the diffusion through the pores of the membrane. The same set of experiments was then repeated with a counter-current flow pattern (Fig. 7b). The results were more erratic but the average values for Ka were similar to those for co-current flow. This result
Fig. 5. Percentage of phenol removed from aqueous solution (3000 ppm) using 1-decanol together with different extractants (5% (v/v)) at 20 ◦ C. Volume ratio (Va :Vo ) = 2:1.
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Fig. 6. Phenol concentration profiles in the (䉬) aqueous and (䊐) organic phases during the extraction of phenol from an aqueous solution with 1-decanol at 20 ◦ C using a membrane contactor. Rea = 4 and Reo = 0.47. Counter-current flow.
Fig. 7. Influence of the organic phase hydrodynamics (Reo ) on the overall mass transfer coefficient defined for the aqueous phase (Ka ) during the extraction of phenol from an aqueous solution with 1-decanol at 20 ◦ C using a membrane contactor (Rea = 4). (a) Reproducible stable values were obtained in co-current flow. (b) Relative to (䊉) co-current flow and (䊐) counter-current flow experiments yielded similar average values, but with a significantly poorer precision.
indicates that for this system the flow pattern does not have a significant influence on the overall mass transfer coefficient, but that the system is more stable when operated with co-current flow patterns. As mentioned in Section 2.3.2, the effect of hydrodynamics during the extraction and stripping processes should be the same.
In all these experiments, the mass balances closed, in general, within 7%. 3.4. Influence of temperature Temperature directly affects the physico-chemical properties of the system (i.e. density, viscosity,
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Fig. 8. Influence of the mole ratio of NaOH to phenol on removal of phenol during stripping of phenol using a membrane contactor at 20 ◦ C. Reo = 0.32 and Res = 7.
interfacial tension and mutual solubility) but not the distribution coefficient. Thus, temperature has an indirect influence on the overall mass transfer coefficient. As the temperature changed from 20 to 40 ◦ C, an increase of 60% in the overall mass transfer coefficient was observed.
3.5. Influence of the mole ratio of NaOH to phenol on stripping efficiency Phenol can be easily stripped from the loaded organic phases by contacting them with an aqueous solution of sodium hydroxide, yielding the phenol sodium
Fig. 9. Phenol concentration ((䉫) aqueous phase; (䊏) organic phase; () stripping phase) as a function of time during a simultaneous extraction and stripping experiment, using 1-decanol as the organic phase and 1.8 M NaOH as the stripping phase at 20 ◦ C. Va = 6 l, Vo = 0.7 l and Vs = 0.4 l. Rea = 7, Reo = 0.32 and Res = 7.
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salt in aqueous solution. Stripping experiments were performed using a membrane contactor at 20 ◦ C with Reo = 0.32 and Res = 7. The experimental protocol was similar to that employed in the extraction experiments described above. The variable to be optimised in this case is the mole ratio of NaOH to phenol. Examination of Fig. 8 shows that it is necessary to utilise a ratio of at least 4 to remove about 97% of the phenol from an organic solution with a phenol concentration of 3 g/l. This ratio is significantly higher than the stoichiometric of the reaction of phenol with NaOH. 3.6. Integrated extraction-stripping process In an integrated extraction-stripping process, the phenol extracted from the aqueous phase into the organic phase, is simultaneously transferred to a stripping solution (aqueous sodium hydroxide). In this system, two membrane contactors are used to carry out the extraction and stripping processes forming a closed loop cycle. This scheme permits continuous regeneration of the organic phase and if the volumetric ratios are properly adjusted, may yield a more concentrated phenol solution. Experiments were carried out at 20 ◦ C using optimum hydrodynamic conditions (co-current flow pattern in both contactors with Rea = 7, Reo = 0.32, and Res = 7). Typical concentration profiles for the case where the aqueous, organic and stripping phase volumes were 6, 0.7, and 0.4 l, respectively, are shown in Fig. 9. A concentration ratio of 14 was obtained. The associated percentage removal of phenol exceeded 99%. Values of phenol recovery and the degree of concentration reached in different experiments are summarised in Table 2.
Table 2 Phenol recovery and concentration ratio reached in several integrated extraction and stripping experiments using membrane contactors at 20 ◦ C (Vo = 0.4 l of 1-decanol in all cases) Va (l)
Vs (l)
Va :Vs ratio
Phenol recovery (%)
Concentration ratio
0.5 0.6 20 10 25
0.6 0.4 0.9 0.4 0.9
8.3 15.0 22.2 25.0 27.8
99.8 99.2 90.3 97.8 93.5
8.2 14 16 21 21
4. Conclusions Experiments were conducted to recover phenol from aqueous solutions using both physical and chemically extractants. Physical extraction using 1-decanol as solvent gave the best results because of its high distribution coefficient and low solubility in water. When membrane contactor experiments were carried out using 1-decanol as solvent, the system was more stable when co-current flows were used. Conditions in which diffusion through the membrane becomes the controlling resistance to the mass transfer were easily reached. A 60% increase in the mass transfer coefficient is observed as the temperature increases from 20 to 40 ◦ C. A concentrated aqueous sodium hydroxide solution is an effective stripping solution. A mole ratio of 4:1 (NaOH:phenol) was needed to recover more than 99% of the extracted phenol. It is possible to carry out extraction and stripping simultaneously, running the system as a closed loop process. Concentration ratios of up to about 20-fold with 98% recovery were obtained.
Acknowledgements This research was funded by CICYT Project QUI1999-0888.
Appendix A A differential mass balance for phenol in the module (Fig. 10) taking into account the transfer of the component from the aqueous into the organic phase is as follows: Qa [ca (x, t) − ca (x + dx, t)] = −Qa dca (x, t) = Ka dAm (ca (x, t) − ca∗ (x, t))
(A.1)
where Qa is the flow rate of the aqueous feed phase, Ka the overall mass transfer coefficient with respect to the aqueous phase, Am the membrane area, ca (x, t) the concentration of phenol in the aqueous solution at the axial position x and at time t, and ca∗ (x, t) is the concentration of phenol in the aqueous phase in equilibrium with that in the organic phase at the same time
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Fig. 10. Schematic of the solute transfer in a membrane contactor system.
and location, co (x, t). The equilibrium relationship can be expressed as ca∗ (x, t) =
co (x, t) D
(A.2)
and taking into account the overall mass balance to the module, Eq. (A.3), Eq. (A.5) can be expressed as Ka Am caout (t) − coout (t)/D (1 + Q) (A.7) = exp − Qa cain (t) − coin (t)/D
where D is the distribution coefficient. The decrease in solute molar flow in the aqueous phase should be equal to the increase in solute molar flow in the organic phase, as expressed in the module overall mass balance:
A simultaneous unsteady-state balance to the aqueous phase feed tank, assuming perfect mixing, gives
Qa (cain (t) − caout (t)) = Qo (coout (t) − coin (t))
At any time the depletion of the amount of solute present in the aqueous phase should be equal to its increase in the organic phase:
(A.3)
where Qo is the flow rate of the organic phase and the superscripts “in” and “out” refer to the module inlet and outlet, respectively (being cain (t) = ca (0, t) and caout (t) = ca (L, t)). A similar mass balance could also be written at any point (x) of the module as follows: Qa (ca (x, t) − caout (t)) = Qo (coout (t) − co (x, t)) (A.4) Substituting co (x, t) from Eq. (A.2) into Eq. (A.4), and ca∗ (x, t) in Eq. (A.1), and integrating this equation between x = 0 (where and ca (x = 0, t) = cain (t)) and x = L (where ca (x = L, t) = caout (t)), the following equation is obtained: caout (t) − coout (t)/D cain (t)(1 + Q) − caout (t)Q − coout /D Ka Am (1 + Q) = exp − Qa Qa Qo D
dcain (t) = Qa [cain (t) − caout (t)] dt
Va (ca0 − cain (t)) = Vo (coin (t) − co0 )
(A.5)
(A.6)
(A.8)
(A.9)
where ca0 = cain (t = 0) and co0 = coin (t = 0) are the initial solute concentration in the aqueous and organic phases, respectively. Combining Eqs. (A.3) and (A.9), the concentration of solute in the organic phase at the module outlet can be expressed as a function of (i) the initial solute concentration in the aforementioned phase, (ii) known quantities, such as phase volumes and flow rates, and (iii) the solute concentration in the aqueous phase: coout (t) =
where Q=
−Va
Qa (cain (t) − caout (t)) Qo Va (ca0 − cain (t)) + + co0 Vo
(A.10)
Combining Eqs. (A.10) and (A.7), allows to obtain an expression for the solute concentration in the aqueous phase at the module outlet [caout (t)] as a function of the inlet concentration [cain (t)] and some
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process parameters: caout (t) =
1 c0 cain (t) − V [ca0 − cain (t)] + o (1 + Q) D Ka Am (1 + Q) exp − Qa 1 co0 in 0 in + Qca (t) + V [ca − ca (t)]+ (1 + Q) D (A.11)
where V =
Va Vo D
(A.12)
Substituting Eq. (A.11) into Eq. (A.8) and integrating between t = 0 and time t: cain (t) =
(ca0 V − co0 /D) (ca0 − co0 /D) + (1 + V ) (1 + V ) Qa (1 + V ) exp − Va (1 + Q) K a Am (1 + Q) t (A.13) × 1 − exp − Qa
Eq. (A.13) is identical to Eqs. (6) and (7) in Section 2.5, when the solute concentration in the aqueous phase is measured in the feed tank [cain (t) = ca ], assuming perfect mixing, and the initial solute concentration in the organic phase, co0 , is zero. Thus, Ka can be obtained from the exponential term, as shown in Eq. (14).
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