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An in situ coupling separation process of electro-electrodialysis with back-extraction
Journal of Membrane Science 255 (2005) 57–65
An in situ coupling separation process of electro-electrodialysis with back-extraction S.S. Yi, Y.C. Lu, G.S. Luo ∗ The State Key Laboratory of Chemical Engineering, Department of Chemical Engineering, Tsinghua University, Beijing 100084, PR China Received 21 September 2004; received in revised form 18 January 2005; accepted 19 January 2005 Available online 21 March 2005
Abstract To intensify mass transfer and simplify the separation processes in recovering and concentrating organic acids from their fermentation broths, an in situ coupling separation process of electro-electrodialysis with back-extraction has been developed. In this new process an oil-in-water disperse system comprised by extractant phase and stripping phase flew in the cathode chamber concurrently. The two steps, including back-extraction and concentration, happen at the same time in one module. The experimental studies on the process were carried out, with a mixed solvent of a kind of tertiary amine with long aliphatic chains and n-octanol as the extractant, and citric acid as the solute. The salt aqueous solution of mono-sodium citrate was used as the stripping phase. The influence of the current density, phase ratio of the extractant phase to the stripping phase, salt concentration of the stripping phase and initial concentration of receptor was investigated. The experimental results showed that the new process integrated the EED process with the back-extraction process in one module successfully, which made the separation process simple and the equipment cost less. The current efficiency of 103.5% and the specific energy consumption of 2.48 kW h/kg were reached. Meanwhile the voltage was not higher than 24 V. Compared with a normal EED process, it is a process with the advantages of higher mass transfer performance and slightly lower energy consumption. © 2005 Elsevier B.V. All rights reserved. Keywords: Separation process; Citric acid; Electro-electrodialysis; Extraction; Integration
1. Introduction A number of carboxylic and hydroxyl-carboxylic acids, such as propionic acid [1], acetic acid [1–3], citric acid [4,5] and lactic acid [6,7], are now almost exclusively produced by microbial fermentation. One of the most important steps in their manufacture process is the recovery of the products from the fermentation broths. The separation process not only accounts for up to 50% of the total production costs [8], but also may create numerous environmental problems. The most commonly applied recovery technique is the acid precipitation [9], which usually yields considerable amounts of solid wastes such as CaSO4 . For example,
∗
Corresponding author. Tel.: +86 10 627 838 70; fax: +86 10 625 627 95. E-mail address: [email protected] (G.S. Luo).
0376-7388/$ – see front matter © 2005 Elsevier B.V. All rights reserved. doi:10.1016/j.memsci.2005.01.020
3–4 t solid wastes are yielded when 1 t citric acid is produced [10]. These large amounts of wastes give rise to serious contamination to the environment. Moreover, the product obtained by this technique is usually at low concentration, e.g. 0.2–1N [11]. Till date new methods for the separation and purification of the fermentation products are highly required. Taking citric acid as an example, we can find that a number of techniques without precipitation have been investigated [4,5,9–22]. Among them electrodialysis (ED) and solvent extraction are two kinds of the most frequently mentioned techniques [9–16]. Voss [12] published an article in 1986 on the deacidification of citric acid solution by ED. Since then, the ED technique has been studied for the recovery of organic acids by many groups. Pinacci and Radaelli [13] studied the effect of cell configuration on the recovery of citric acid, and presented a suitable compart-
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ment configuration constructed by cation-exchange membranes and bipolar membranes. Xu and Yang [10] investigated the production of citric acid from sodium citrate by using bipolar membrane, and the specific energy consumption was about 3–8 kW h/kg. Reactive liquid–liquid extraction by an appropriate extractant was found to be a promising alternative to the conventional technique, and tertiary amines with long aliphatic chains (TAA) were suggested as the high capacity complexants [9]. This technique with sodium hydroxide as stripping agent was applied in industry by Pfizer Inc. in Europe and Haarman and Reimer Corp. in the USA [15,16]. However, in these circumstances, the catholyte can easily form emulsions. This impairs the subsequent phase separation, an operation that can eventually fail in the worst cases. In addition, it produces low concentration sodium citrate aqueous solution in the backextraction process, which is required to be transferred to the final product, citric acid with ED. The whole process is a combination process of extraction with electrodialysis. In previous work, we found that the extraction process can be intensified by an external electric field. The techniques of two-phase electrophoresis (TPE) and two-phase electroelectrodialysis (TPEED) were suggested to recover organic acids from their dilute solutions [23–26]. In the TPE process an electric field was applied to the two-phase system directly. The distribution coefficient of organic acids was successfully increased when n-butanol was used as the extractant [23]. But due to very low electric conductivity of the organic phase and the direct contact of the two phases, the energy consumption was very high and the interphase was unstable. In 2004, Luo et al. [25,26] proposed a new technique called twophase electro-electrodialysis (TPEED), which was a coupling process of TPE with electro-electrodialysis (EED). In the TPEED process the two phases were separated by a piece of anion exchange membrane, and pure organic solvent or W/O emulsion was used as the extractant. Using the new technique to separate and concentrate organic acids, it can overcome the electro-osmosis and osmosis of water, and also the back diffusion can be well controlled. So the higher recovery ratio and higher concentrated ratio were reached. However, the specific energy consumption was still as high as about 8–10 kW h/kg citric acid. Furthermore the selection of the organic solvent is difficult, because the organic phase is required to have good electric conductivity in the TPE or TPEED process. In order to develop a compact, safe, energy-saving, and environment-friendly process for producing organic acids from their fermentation broths, a modified TPEED technique is developed in this paper. It is an in situ coupling separation process of electro-electrodialysis with back-extraction, i.e., the back-extraction and the concentration process are integrated in one unit operation. The objective of this work is to study the performance of the modified TPEED technique with citric acid as a model product, and the citrate aqueous solution as the stripping agent. The influence of dif-
ferent operating conditions on the process was investigated, and the comparison of the process with EED was carried out.
2. Experimental 2.1. Reagents and anion-exchange membrane Citric acid and sodium hydroxide were purchased from Beijing Fine Chemical reagent Ltd. and n-octanol was obtained from Tianjin Bodi Chemical Company Ltd. They were all of AR grade, and used without further purification. A kind of tertiary amine with C8–C10 aliphatic chains named 7301 was kindly provided by Shuzhou Organic Solvent Company. JIPE-AM-203 anion-exchange membrane with polyethylene as substrate, RN+ (CH3 )3 as functional group, exchange capacity of 1.9 mequiv./g (dry membrane), water content of about 25%, resistance of about 10 cm2 , thickness of about 0.3 mm and selectivity of 96% was purchased from Shanghai Qiulong Chemical Company, Ltd. The mono-sodium citrate aqueous solutions were prepared in our laboratory. 2.2. The experimental set-up and procedure Fig. 1 is a schematic diagram of the experimental setup employed in this work. The system was comprised of an electro-electrodialyzer, four storage tanks, two peristaltic pumps and an adjustable dc power supply. The electroelectrodialyzer made by PTFE was divided into an anode and a cathode compartment by an anion-exchange membrane. The thicknesses of the two compartments were 5 mm, respectively. The effective membrane area was 60 cm2 . Two electrodes connected to the dc power supply were made of platinum. The constant current density in the electroelectrodialysis process was controlled by the dc power supply. Two of the storage tanks and two pumps were used to
Fig. 1. The set-up of experimental apparatus. (1) Disperse system feed tank, (2) receptor feed tank, (3) disperse system collection tank, (4) receptor collection tank, (5) catholyte pump, (6) anolyte pump, (7) cathode chamber, (8) anode chamber, (9) anion-exchange membrane, (10) magnetic stirrer.
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continuously deliver anolyte and catholyte into the electroelectrodialyzer with the flow rate of about 0.0002 m3 /h, respectively. The other two tanks were used to collect the efflux of anolyte and catholyte, respectively. Correspondingly, the residence time of electrolyte was about 540 s. A disperse system by dispersing the loading extractant into the mono-sodium citrate aqueous solutions was taken as catholyte. A magnetic stirrer was used to maintain the disperse system stable. The whole set-up was placed in environment at room temperature without additional temperature controller. The extractant used in this work was 20–40 vol.% 7301 solution, n-octanol as solvent. After extracting citric acid from a citric acid solution, the loading extractant was mixed with a mono-sodium citrate aqueous solution to form disperse system, which was used as catholyte. A desired concentration solution of citric acid was applied as the concentrated phase. When the steady state operation at different conditions reached, the samples of receptor were collected and the citric acid concentration was analyzed by titration. To simplify the data analysis, we suppose that only the mono charged species, H2 A− passes through the membrane. It approaches to the true case as the residence time is short to reach high recovery ratio in processing. And it is almost satisfied in the paper. Furthermore, five parameters to evaluate the modified technique, including solute flux J, recovery ratio E, current efficiency η, average cell voltage U and specific energy consumption Wt are obtained as follows: J=
CO1 vO1 − CI1 vI1 A
E=
CO1 vO1 − CI1 vI1 CI2 vI2 φ
η (%) =
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CO1 vO1 − CI1 vI1 · zF · 100 IA
t A · 0 IU dt Wt = (CO1 vO1 − CI1 vI1 ) · M · t where t is the time, ϕ the volume fraction of loading extractant in the disperse system, z the mean number of negative charges carried by the species crossing the membrane, F the Faraday constant, I the applied current density, A the effective membrane area, U the cell voltage, and M the molecular weight of citric acid. CO1 and CI1 are the citric acid molarities of analyte at outlet and inlet respectively, CI2 the citric acid molarities of the loading extractant at inlet. vO1 and vI1 are the volume flow rates of anolyte at outlet and inlet respectively, and vI2 is the volume flow rate of the disperse system at inlet. 2.3. Mechanism of the coupling separation process The in situ coupling process involves two main mass transfer steps, the migration of citric acid from extractant phase to stripping phase, and the transfer of citrate anions from stripping phase to receptor by passing through the anion-exchange membrane under an electric field, as shown in Fig. 2. Herein we denote the molecule of citric acid as H3 A, and the monosodium citrate in stripping solution as H2 A− . Citric acid is inert to electrode reactions. Therefore the cathode reaction is 2H+ + 2e → H2 ↑ or 2H2 O + 2e → 2OH− + H2 ↑, and the anode reaction is H2 O → 2H+ + 21 O2 ↑ + 2e. The whole reaction can be regarded as a splitting water process. There exist two possible routes for the citric acid migration from extractant phase to stripping phase. One is that the OH− reacts with H3 A directly to form H2 A− and enter stripping phase. The other is that the OH− reacts with the H2 A− in
Fig. 2. A possible mass transfer mechanism in the coupling separation process. (1) OH−1 , (2) H2 A− or HA2− , (3) droplets of solvent, (4) H3 A, (5) H+ .
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stripping phase to form HA2− firstly, and the HA2− in sequence reacts with the H3 A to form H2 A− and enter stripping phase. Thus, the migration of H2 A− , HA2− and OH− toward anode chamber may all proceed in the stripping phase. In experiments, because the H2 A− was rich and of absolutely large amount in stripping phase, and the cathode chamber was somewhat thick, the H2 A− occupied the majority in what passed through the membrane and the z approximates 1. After entering the anode chamber, it combines with H+ generated by the anode reaction to form H3 A.
3. Results and discussions 3.1. Extraction behavior of the mixed solvent to citric acid It is important to select an appropriate extractant by some special criterions, such as high extraction coefficient, selectivity and capacity, as well as short split-phase time. Based on the literatures’ results [9,16,27–28], the mixed solvents of 7301 and n-octanol, with the 7301 volume fraction of 20, 30 and 40% were used as the extractants. The distribution coefficients at 25 ◦ C are illustrated in Fig. 3. It decreases from far over 1 to 0.2 rapidly as the concentration of citric acid in aqueous solutions changes from 0.04 to 0.5 mol/L, and then slightly decreases with the further increase of the concentration. The distribution coefficient is positive correlated to the volume fraction of 7301 in the extractant, and so is
Fig. 3. Distribution coefficient of citric acid in the mixed solvent (T = 25 ◦ C)
the viscosity of the extractant. Since high viscosity may result in difficulty in phase separation, the mixed solvent of 30 vol.% 7301 in n-octanol was used in following experiments. 3.2. Influence of current density on the modified process The mass transfer performance of the new process is strongly dependent on current density, volume phase ratio of extractant phase to stripping phase, salt concentration of stripping phase and citric acid concentration of receptor. The influence of these factors on the mass transfer performance was investigated. Fig. 4 shows the influence of current density on the solute flux, recovery ratio, overall current
Fig. 4. Influence of current density on the separation process (cW = 0.73 mol/L, cO = 0.56 mol/L, cS = 0.72 mol/L, r = 1).
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efficiency and average cell voltage. An increase in current density resulted in the increasing of solute flux, recovery ratio and current efficiency. This is similar to that of most of EED processes operated under lower current density [22,25]. The solute flux reached as much as 0.92 kg/m2 h, much greater than 0.25 kg/m2 h in a normal ED process [22], which means the mass transfer enhanced greatly. Also the current efficiency was larger than 0.60, larger than 0.52 in a normal ED process [22]. Recovery ratio over 60% was reached in a single pass. It shows that citric acid was back-extracted from the extractant phase and concentrated into the receptor successfully. The applied current density in the modified process could be 180 A/m2 or even higher, and the corresponding average cell voltage was about 20 V, far less than that of TPEED process [25]. This was because the high salt concentration of the stripping phase in disperse system improved electric conducting in the modified process. With the increasing of current density, the electric power consumption increased from 9.6 to 23.6 W, in which about 85% transferred to Joule heat (the rest was for water electrolysis). By estimation, although the heat could dissipate from the wall of electrolyzer partly, the temperature of electrolyte might increase about 15–30 ◦ C. The increasing of temperature may improve the electric conducting in the system, which is profit for compressing energy consumption in higher current density. Besides, higher temperature may strengthen heat turbulence leading to lower current efficiency. In the paper, as changing any single parameter except for the current density, the fluctuation of temperature was relative little, maybe less than 5 ◦ C, and its effect on the revealed characters is quite little. As regards to
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the current density, the down curved line of cell voltage variation is partly reflected with the influence of processing temperature. Meanwhile, the increasing of current efficiency shown in Fig. 4(c) is slower than that under isothermal condition. 3.3. Influence of phase ratio on the modified process Fig. 5 shows the influence of the phase ratio in the disperse system on the modified TPEED process. It tells us that the solute flux and the current efficiency were increased with the phase ratio increasing, but the recovery ratio and the average cell voltage were decreased. When the phase ratio was 5, the solute flux was larger than 0.82 kg/m2 h, the current efficiency was near 0.7, and the average cell voltage was about 16 V. These results mean that the conductivity of the disperse system was increased with the phase ratio increasing. In the experiments, we observed that the phase separation process at high phase ratio was slower than that at low phase ratio. It seems that the better stability of the disperse system resulted in the higher electric conductivity. 3.4. Influence of salt concentration of the stripping phase on the modified process The increase of the salt concentration in the stripping phase could lead to the growing of the electric conductivity of the system. Therefore the average cell voltage was decreased slightly with an increase of the salt concentration. Meanwhile the solute flux, recovery ratio and current efficiency were increased slowly, as shown in Fig. 6.
Fig. 5. Influence of phase ratio on the separation process (cW = 0.70 mol/L, cO = 0.53 mol/L, cS = 0.98 mol/L, I = 150 A/m2 ).
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Fig. 6. Influence of salt concentration in the stripping phase on the separation process (cW = 1.11 mol/L, cO = 0.51 mol/L, r = 2, I = 150 A/m2 ).
3.5. Influence of the initial citric acid concentration of receptor on the modified process The initial concentration of receptor was another factor to influence the modified process. As shown in Fig. 7, there was a maximum value for the solute flux, recovery ratio and current efficiency when the initial concentration of receptor was changed. But the average cell voltage was nearly kept
at a constant value. As the initial concentration was higher than 1.35 mol/L, the solute flux, recovery ratio and current efficiency began to decrease. For the upward region, the inhibition to proton leaking may be a possible cause, i.e., the electro-migration of H+ in anode chamber will slow down for the concentration of both H+ and acid group ion increasing in the case. For the downward region, the possible reason was the dominant back diffusion.
Fig. 7. Influence of the initial citric acid concentration in receptor (cO = 0.42 mol/L, cS = 0.98 mol/L, r = 2, I = 150 A/m2 ).
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Fig. 8. Effect of operation conditions on the specific energy consumption. (1) cW = 0.7301 mol/L, cO = 0.5610 mol/L, cS = 0.7187 mol/L, r = 1. (2) cW = 0.6995 mol/L, cO = 0.5258 mol/L, cS = 0.9822 mol/L, I = 150 A/m2 . (3) cW = 1.1055 mol/L, cO = 0.5148 mol/L, r = 2, I = 150 A/m2 . (4) cO = 0.4163 mol/L, cS = 0.9822 mol/L, r = 2, I = 150 A/m2 .
3.6. Influence of operation conditions on the specific energy consumptions
value was 8.86 kW h/kg. These are less than that in the TPE or TPEED process [24,25].
A separation process not only must be technically feasible, but also should be less expensive. To evaluate the economical feasibility of the modified process, the specific energy consumptions under different operation conditions were calculated and the results are shown in Fig. 8. The specific energy consumptions decreased as the current density, phase ratio, salt concentration or initial citric acid concentration was increased. But once the initial citric acid concentration of receptor was larger than 1.35 mol/L, the specific energy consumption was increased. So the major factors to the specific energy consumption were the phase ratio and the initial concentration of receptor. The minimum value of the specific energy consumption was 2.48 kW h/kg, and the maximum
3.7. Comparison of the modified TPEED process with EED process A normal EED process in the same experimental set-up with the same anion-exchange membrane has been carried out. The experimental conditions and the compared results are listed in Table 1. The performance of the modified TPEED process was similar to that of the normal EED process. For example, the voltages of TPEED were as low as that of EED at the same current density, and the current efficiency was close with each other. The results of comparison show the new process has the advantages of simpleness, low energy consumption, high mass flux and high current efficiency.
Table 1 Comparison of TPEED with EED for concentrating citric acid solution Initial citric acid concentration, cW (mol/L) Salt concentration, cS (mol/L) Method Volume phase ratio, r Recovery ratio, E Concentrated ratio, R Current efficiency, η Solute flux, J (mol/m2 min) Specific energy consumption
0.70 0.98 TPEED 2 0.34 1.08 0.67 0.06 3.85
0.70 0.98 TPEED 5 0.32 1.10 0.76 0.07 2.89
0.68
0.19 1.07 0.79 0.07 3.17
Wt (kW h/kg citric acid) Average cell voltage, U (V) Current density, I (A/m2 )
20.1 150
17.1 150
19.3 150
EED
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4. Conclusions In order to make the separation process compact, simple and high efficiency, a coupling process of EED with backextraction for recovering and concentrating organic acids was developed based on TPEED process. The performance of the modified process has been studied with the mixed organic solvent of 7301 and n-octanol as the extractant. The experimental results show that it is an effective technique to recover citric acid. At suitable operation condition, a recovery ratio more than 60% could be reached in a single pass. The average cell voltage could be obviously decreased, and the specific energy consumption was limited and satisfied. And also our experimental results shows that the major factors to influence the performance of the process is the current density, phase ratio of the extractant phase to the stripping phase, the salt concentration of the stripping solution, and the initial citric acid concentration in the receptor. The increase in these experimental factors resulted in an increase of the current efficiency and mass transfer flux. But as the initial citric acid concentration was larger than 1.35 mol/L, the mass flux and current efficiency decreased. In general, an in situ coupling separation process of electro-electrodialysis with back-extraction is developed. This new technique could provide a simple separation process, higher current efficiency and recovery ratio, and lower average cell voltage and specific energy consumption. Acknowledgment We wish to acknowledge the support of the National Science Foundation of China and National Basic Research Program of China (2003CB615701) on this work gratefully.
Nomenclature A cW cO cS D E I J M r U v Wt
effective membrane area (m2 ) concentration of citric acid in aqueous phase (mol/L) concentration of citric acid in organic phase (mol/L) concentration of citrate in stripping phase (mol/L) distribution coefficient recovery ratio current density (A/m2 ) solute flux of citric acid (mol/m2 min) molecular weight of citric acid (kg/mol) volume phase ratio of extractant phase to stripping phase cell voltage (V) volume flow rate (m3 /h) specific energy consumption per unit mass (kW h/kg)
Greek letters η current efficiency φ volume fraction of extractant phase in the disperse system
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An in situ coupling separation process of electro-electrodialysis with back-extraction S.S. Yi, Y.C. Lu, G.S. Luo ∗ The State Key Laboratory of Chemical Engineering, Department of Chemical Engineering, Tsinghua University, Beijing 100084, PR China Received 21 September 2004; received in revised form 18 January 2005; accepted 19 January 2005 Available online 21 March 2005
Abstract To intensify mass transfer and simplify the separation processes in recovering and concentrating organic acids from their fermentation broths, an in situ coupling separation process of electro-electrodialysis with back-extraction has been developed. In this new process an oil-in-water disperse system comprised by extractant phase and stripping phase flew in the cathode chamber concurrently. The two steps, including back-extraction and concentration, happen at the same time in one module. The experimental studies on the process were carried out, with a mixed solvent of a kind of tertiary amine with long aliphatic chains and n-octanol as the extractant, and citric acid as the solute. The salt aqueous solution of mono-sodium citrate was used as the stripping phase. The influence of the current density, phase ratio of the extractant phase to the stripping phase, salt concentration of the stripping phase and initial concentration of receptor was investigated. The experimental results showed that the new process integrated the EED process with the back-extraction process in one module successfully, which made the separation process simple and the equipment cost less. The current efficiency of 103.5% and the specific energy consumption of 2.48 kW h/kg were reached. Meanwhile the voltage was not higher than 24 V. Compared with a normal EED process, it is a process with the advantages of higher mass transfer performance and slightly lower energy consumption. © 2005 Elsevier B.V. All rights reserved. Keywords: Separation process; Citric acid; Electro-electrodialysis; Extraction; Integration
1. Introduction A number of carboxylic and hydroxyl-carboxylic acids, such as propionic acid [1], acetic acid [1–3], citric acid [4,5] and lactic acid [6,7], are now almost exclusively produced by microbial fermentation. One of the most important steps in their manufacture process is the recovery of the products from the fermentation broths. The separation process not only accounts for up to 50% of the total production costs [8], but also may create numerous environmental problems. The most commonly applied recovery technique is the acid precipitation [9], which usually yields considerable amounts of solid wastes such as CaSO4 . For example,
∗
Corresponding author. Tel.: +86 10 627 838 70; fax: +86 10 625 627 95. E-mail address: [email protected] (G.S. Luo).
0376-7388/$ – see front matter © 2005 Elsevier B.V. All rights reserved. doi:10.1016/j.memsci.2005.01.020
3–4 t solid wastes are yielded when 1 t citric acid is produced [10]. These large amounts of wastes give rise to serious contamination to the environment. Moreover, the product obtained by this technique is usually at low concentration, e.g. 0.2–1N [11]. Till date new methods for the separation and purification of the fermentation products are highly required. Taking citric acid as an example, we can find that a number of techniques without precipitation have been investigated [4,5,9–22]. Among them electrodialysis (ED) and solvent extraction are two kinds of the most frequently mentioned techniques [9–16]. Voss [12] published an article in 1986 on the deacidification of citric acid solution by ED. Since then, the ED technique has been studied for the recovery of organic acids by many groups. Pinacci and Radaelli [13] studied the effect of cell configuration on the recovery of citric acid, and presented a suitable compart-
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ment configuration constructed by cation-exchange membranes and bipolar membranes. Xu and Yang [10] investigated the production of citric acid from sodium citrate by using bipolar membrane, and the specific energy consumption was about 3–8 kW h/kg. Reactive liquid–liquid extraction by an appropriate extractant was found to be a promising alternative to the conventional technique, and tertiary amines with long aliphatic chains (TAA) were suggested as the high capacity complexants [9]. This technique with sodium hydroxide as stripping agent was applied in industry by Pfizer Inc. in Europe and Haarman and Reimer Corp. in the USA [15,16]. However, in these circumstances, the catholyte can easily form emulsions. This impairs the subsequent phase separation, an operation that can eventually fail in the worst cases. In addition, it produces low concentration sodium citrate aqueous solution in the backextraction process, which is required to be transferred to the final product, citric acid with ED. The whole process is a combination process of extraction with electrodialysis. In previous work, we found that the extraction process can be intensified by an external electric field. The techniques of two-phase electrophoresis (TPE) and two-phase electroelectrodialysis (TPEED) were suggested to recover organic acids from their dilute solutions [23–26]. In the TPE process an electric field was applied to the two-phase system directly. The distribution coefficient of organic acids was successfully increased when n-butanol was used as the extractant [23]. But due to very low electric conductivity of the organic phase and the direct contact of the two phases, the energy consumption was very high and the interphase was unstable. In 2004, Luo et al. [25,26] proposed a new technique called twophase electro-electrodialysis (TPEED), which was a coupling process of TPE with electro-electrodialysis (EED). In the TPEED process the two phases were separated by a piece of anion exchange membrane, and pure organic solvent or W/O emulsion was used as the extractant. Using the new technique to separate and concentrate organic acids, it can overcome the electro-osmosis and osmosis of water, and also the back diffusion can be well controlled. So the higher recovery ratio and higher concentrated ratio were reached. However, the specific energy consumption was still as high as about 8–10 kW h/kg citric acid. Furthermore the selection of the organic solvent is difficult, because the organic phase is required to have good electric conductivity in the TPE or TPEED process. In order to develop a compact, safe, energy-saving, and environment-friendly process for producing organic acids from their fermentation broths, a modified TPEED technique is developed in this paper. It is an in situ coupling separation process of electro-electrodialysis with back-extraction, i.e., the back-extraction and the concentration process are integrated in one unit operation. The objective of this work is to study the performance of the modified TPEED technique with citric acid as a model product, and the citrate aqueous solution as the stripping agent. The influence of dif-
ferent operating conditions on the process was investigated, and the comparison of the process with EED was carried out.
2. Experimental 2.1. Reagents and anion-exchange membrane Citric acid and sodium hydroxide were purchased from Beijing Fine Chemical reagent Ltd. and n-octanol was obtained from Tianjin Bodi Chemical Company Ltd. They were all of AR grade, and used without further purification. A kind of tertiary amine with C8–C10 aliphatic chains named 7301 was kindly provided by Shuzhou Organic Solvent Company. JIPE-AM-203 anion-exchange membrane with polyethylene as substrate, RN+ (CH3 )3 as functional group, exchange capacity of 1.9 mequiv./g (dry membrane), water content of about 25%, resistance of about 10 cm2 , thickness of about 0.3 mm and selectivity of 96% was purchased from Shanghai Qiulong Chemical Company, Ltd. The mono-sodium citrate aqueous solutions were prepared in our laboratory. 2.2. The experimental set-up and procedure Fig. 1 is a schematic diagram of the experimental setup employed in this work. The system was comprised of an electro-electrodialyzer, four storage tanks, two peristaltic pumps and an adjustable dc power supply. The electroelectrodialyzer made by PTFE was divided into an anode and a cathode compartment by an anion-exchange membrane. The thicknesses of the two compartments were 5 mm, respectively. The effective membrane area was 60 cm2 . Two electrodes connected to the dc power supply were made of platinum. The constant current density in the electroelectrodialysis process was controlled by the dc power supply. Two of the storage tanks and two pumps were used to
Fig. 1. The set-up of experimental apparatus. (1) Disperse system feed tank, (2) receptor feed tank, (3) disperse system collection tank, (4) receptor collection tank, (5) catholyte pump, (6) anolyte pump, (7) cathode chamber, (8) anode chamber, (9) anion-exchange membrane, (10) magnetic stirrer.
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continuously deliver anolyte and catholyte into the electroelectrodialyzer with the flow rate of about 0.0002 m3 /h, respectively. The other two tanks were used to collect the efflux of anolyte and catholyte, respectively. Correspondingly, the residence time of electrolyte was about 540 s. A disperse system by dispersing the loading extractant into the mono-sodium citrate aqueous solutions was taken as catholyte. A magnetic stirrer was used to maintain the disperse system stable. The whole set-up was placed in environment at room temperature without additional temperature controller. The extractant used in this work was 20–40 vol.% 7301 solution, n-octanol as solvent. After extracting citric acid from a citric acid solution, the loading extractant was mixed with a mono-sodium citrate aqueous solution to form disperse system, which was used as catholyte. A desired concentration solution of citric acid was applied as the concentrated phase. When the steady state operation at different conditions reached, the samples of receptor were collected and the citric acid concentration was analyzed by titration. To simplify the data analysis, we suppose that only the mono charged species, H2 A− passes through the membrane. It approaches to the true case as the residence time is short to reach high recovery ratio in processing. And it is almost satisfied in the paper. Furthermore, five parameters to evaluate the modified technique, including solute flux J, recovery ratio E, current efficiency η, average cell voltage U and specific energy consumption Wt are obtained as follows: J=
CO1 vO1 − CI1 vI1 A
E=
CO1 vO1 − CI1 vI1 CI2 vI2 φ
η (%) =
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CO1 vO1 − CI1 vI1 · zF · 100 IA
t A · 0 IU dt Wt = (CO1 vO1 − CI1 vI1 ) · M · t where t is the time, ϕ the volume fraction of loading extractant in the disperse system, z the mean number of negative charges carried by the species crossing the membrane, F the Faraday constant, I the applied current density, A the effective membrane area, U the cell voltage, and M the molecular weight of citric acid. CO1 and CI1 are the citric acid molarities of analyte at outlet and inlet respectively, CI2 the citric acid molarities of the loading extractant at inlet. vO1 and vI1 are the volume flow rates of anolyte at outlet and inlet respectively, and vI2 is the volume flow rate of the disperse system at inlet. 2.3. Mechanism of the coupling separation process The in situ coupling process involves two main mass transfer steps, the migration of citric acid from extractant phase to stripping phase, and the transfer of citrate anions from stripping phase to receptor by passing through the anion-exchange membrane under an electric field, as shown in Fig. 2. Herein we denote the molecule of citric acid as H3 A, and the monosodium citrate in stripping solution as H2 A− . Citric acid is inert to electrode reactions. Therefore the cathode reaction is 2H+ + 2e → H2 ↑ or 2H2 O + 2e → 2OH− + H2 ↑, and the anode reaction is H2 O → 2H+ + 21 O2 ↑ + 2e. The whole reaction can be regarded as a splitting water process. There exist two possible routes for the citric acid migration from extractant phase to stripping phase. One is that the OH− reacts with H3 A directly to form H2 A− and enter stripping phase. The other is that the OH− reacts with the H2 A− in
Fig. 2. A possible mass transfer mechanism in the coupling separation process. (1) OH−1 , (2) H2 A− or HA2− , (3) droplets of solvent, (4) H3 A, (5) H+ .
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stripping phase to form HA2− firstly, and the HA2− in sequence reacts with the H3 A to form H2 A− and enter stripping phase. Thus, the migration of H2 A− , HA2− and OH− toward anode chamber may all proceed in the stripping phase. In experiments, because the H2 A− was rich and of absolutely large amount in stripping phase, and the cathode chamber was somewhat thick, the H2 A− occupied the majority in what passed through the membrane and the z approximates 1. After entering the anode chamber, it combines with H+ generated by the anode reaction to form H3 A.
3. Results and discussions 3.1. Extraction behavior of the mixed solvent to citric acid It is important to select an appropriate extractant by some special criterions, such as high extraction coefficient, selectivity and capacity, as well as short split-phase time. Based on the literatures’ results [9,16,27–28], the mixed solvents of 7301 and n-octanol, with the 7301 volume fraction of 20, 30 and 40% were used as the extractants. The distribution coefficients at 25 ◦ C are illustrated in Fig. 3. It decreases from far over 1 to 0.2 rapidly as the concentration of citric acid in aqueous solutions changes from 0.04 to 0.5 mol/L, and then slightly decreases with the further increase of the concentration. The distribution coefficient is positive correlated to the volume fraction of 7301 in the extractant, and so is
Fig. 3. Distribution coefficient of citric acid in the mixed solvent (T = 25 ◦ C)
the viscosity of the extractant. Since high viscosity may result in difficulty in phase separation, the mixed solvent of 30 vol.% 7301 in n-octanol was used in following experiments. 3.2. Influence of current density on the modified process The mass transfer performance of the new process is strongly dependent on current density, volume phase ratio of extractant phase to stripping phase, salt concentration of stripping phase and citric acid concentration of receptor. The influence of these factors on the mass transfer performance was investigated. Fig. 4 shows the influence of current density on the solute flux, recovery ratio, overall current
Fig. 4. Influence of current density on the separation process (cW = 0.73 mol/L, cO = 0.56 mol/L, cS = 0.72 mol/L, r = 1).
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efficiency and average cell voltage. An increase in current density resulted in the increasing of solute flux, recovery ratio and current efficiency. This is similar to that of most of EED processes operated under lower current density [22,25]. The solute flux reached as much as 0.92 kg/m2 h, much greater than 0.25 kg/m2 h in a normal ED process [22], which means the mass transfer enhanced greatly. Also the current efficiency was larger than 0.60, larger than 0.52 in a normal ED process [22]. Recovery ratio over 60% was reached in a single pass. It shows that citric acid was back-extracted from the extractant phase and concentrated into the receptor successfully. The applied current density in the modified process could be 180 A/m2 or even higher, and the corresponding average cell voltage was about 20 V, far less than that of TPEED process [25]. This was because the high salt concentration of the stripping phase in disperse system improved electric conducting in the modified process. With the increasing of current density, the electric power consumption increased from 9.6 to 23.6 W, in which about 85% transferred to Joule heat (the rest was for water electrolysis). By estimation, although the heat could dissipate from the wall of electrolyzer partly, the temperature of electrolyte might increase about 15–30 ◦ C. The increasing of temperature may improve the electric conducting in the system, which is profit for compressing energy consumption in higher current density. Besides, higher temperature may strengthen heat turbulence leading to lower current efficiency. In the paper, as changing any single parameter except for the current density, the fluctuation of temperature was relative little, maybe less than 5 ◦ C, and its effect on the revealed characters is quite little. As regards to
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the current density, the down curved line of cell voltage variation is partly reflected with the influence of processing temperature. Meanwhile, the increasing of current efficiency shown in Fig. 4(c) is slower than that under isothermal condition. 3.3. Influence of phase ratio on the modified process Fig. 5 shows the influence of the phase ratio in the disperse system on the modified TPEED process. It tells us that the solute flux and the current efficiency were increased with the phase ratio increasing, but the recovery ratio and the average cell voltage were decreased. When the phase ratio was 5, the solute flux was larger than 0.82 kg/m2 h, the current efficiency was near 0.7, and the average cell voltage was about 16 V. These results mean that the conductivity of the disperse system was increased with the phase ratio increasing. In the experiments, we observed that the phase separation process at high phase ratio was slower than that at low phase ratio. It seems that the better stability of the disperse system resulted in the higher electric conductivity. 3.4. Influence of salt concentration of the stripping phase on the modified process The increase of the salt concentration in the stripping phase could lead to the growing of the electric conductivity of the system. Therefore the average cell voltage was decreased slightly with an increase of the salt concentration. Meanwhile the solute flux, recovery ratio and current efficiency were increased slowly, as shown in Fig. 6.
Fig. 5. Influence of phase ratio on the separation process (cW = 0.70 mol/L, cO = 0.53 mol/L, cS = 0.98 mol/L, I = 150 A/m2 ).
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Fig. 6. Influence of salt concentration in the stripping phase on the separation process (cW = 1.11 mol/L, cO = 0.51 mol/L, r = 2, I = 150 A/m2 ).
3.5. Influence of the initial citric acid concentration of receptor on the modified process The initial concentration of receptor was another factor to influence the modified process. As shown in Fig. 7, there was a maximum value for the solute flux, recovery ratio and current efficiency when the initial concentration of receptor was changed. But the average cell voltage was nearly kept
at a constant value. As the initial concentration was higher than 1.35 mol/L, the solute flux, recovery ratio and current efficiency began to decrease. For the upward region, the inhibition to proton leaking may be a possible cause, i.e., the electro-migration of H+ in anode chamber will slow down for the concentration of both H+ and acid group ion increasing in the case. For the downward region, the possible reason was the dominant back diffusion.
Fig. 7. Influence of the initial citric acid concentration in receptor (cO = 0.42 mol/L, cS = 0.98 mol/L, r = 2, I = 150 A/m2 ).
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Fig. 8. Effect of operation conditions on the specific energy consumption. (1) cW = 0.7301 mol/L, cO = 0.5610 mol/L, cS = 0.7187 mol/L, r = 1. (2) cW = 0.6995 mol/L, cO = 0.5258 mol/L, cS = 0.9822 mol/L, I = 150 A/m2 . (3) cW = 1.1055 mol/L, cO = 0.5148 mol/L, r = 2, I = 150 A/m2 . (4) cO = 0.4163 mol/L, cS = 0.9822 mol/L, r = 2, I = 150 A/m2 .
3.6. Influence of operation conditions on the specific energy consumptions
value was 8.86 kW h/kg. These are less than that in the TPE or TPEED process [24,25].
A separation process not only must be technically feasible, but also should be less expensive. To evaluate the economical feasibility of the modified process, the specific energy consumptions under different operation conditions were calculated and the results are shown in Fig. 8. The specific energy consumptions decreased as the current density, phase ratio, salt concentration or initial citric acid concentration was increased. But once the initial citric acid concentration of receptor was larger than 1.35 mol/L, the specific energy consumption was increased. So the major factors to the specific energy consumption were the phase ratio and the initial concentration of receptor. The minimum value of the specific energy consumption was 2.48 kW h/kg, and the maximum
3.7. Comparison of the modified TPEED process with EED process A normal EED process in the same experimental set-up with the same anion-exchange membrane has been carried out. The experimental conditions and the compared results are listed in Table 1. The performance of the modified TPEED process was similar to that of the normal EED process. For example, the voltages of TPEED were as low as that of EED at the same current density, and the current efficiency was close with each other. The results of comparison show the new process has the advantages of simpleness, low energy consumption, high mass flux and high current efficiency.
Table 1 Comparison of TPEED with EED for concentrating citric acid solution Initial citric acid concentration, cW (mol/L) Salt concentration, cS (mol/L) Method Volume phase ratio, r Recovery ratio, E Concentrated ratio, R Current efficiency, η Solute flux, J (mol/m2 min) Specific energy consumption
0.70 0.98 TPEED 2 0.34 1.08 0.67 0.06 3.85
0.70 0.98 TPEED 5 0.32 1.10 0.76 0.07 2.89
0.68
0.19 1.07 0.79 0.07 3.17
Wt (kW h/kg citric acid) Average cell voltage, U (V) Current density, I (A/m2 )
20.1 150
17.1 150
19.3 150
EED
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4. Conclusions In order to make the separation process compact, simple and high efficiency, a coupling process of EED with backextraction for recovering and concentrating organic acids was developed based on TPEED process. The performance of the modified process has been studied with the mixed organic solvent of 7301 and n-octanol as the extractant. The experimental results show that it is an effective technique to recover citric acid. At suitable operation condition, a recovery ratio more than 60% could be reached in a single pass. The average cell voltage could be obviously decreased, and the specific energy consumption was limited and satisfied. And also our experimental results shows that the major factors to influence the performance of the process is the current density, phase ratio of the extractant phase to the stripping phase, the salt concentration of the stripping solution, and the initial citric acid concentration in the receptor. The increase in these experimental factors resulted in an increase of the current efficiency and mass transfer flux. But as the initial citric acid concentration was larger than 1.35 mol/L, the mass flux and current efficiency decreased. In general, an in situ coupling separation process of electro-electrodialysis with back-extraction is developed. This new technique could provide a simple separation process, higher current efficiency and recovery ratio, and lower average cell voltage and specific energy consumption. Acknowledgment We wish to acknowledge the support of the National Science Foundation of China and National Basic Research Program of China (2003CB615701) on this work gratefully.
Nomenclature A cW cO cS D E I J M r U v Wt
effective membrane area (m2 ) concentration of citric acid in aqueous phase (mol/L) concentration of citric acid in organic phase (mol/L) concentration of citrate in stripping phase (mol/L) distribution coefficient recovery ratio current density (A/m2 ) solute flux of citric acid (mol/m2 min) molecular weight of citric acid (kg/mol) volume phase ratio of extractant phase to stripping phase cell voltage (V) volume flow rate (m3 /h) specific energy consumption per unit mass (kW h/kg)
Greek letters η current efficiency φ volume fraction of extractant phase in the disperse system
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