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J. of Supercritical Fluids 25 (2003) 45 /55 www.elsevier.com/locate/supflu

Supercritical fluid extraction of ethanol from aqueous solutions M. Budich 1, G. Brunner Technical University of Hamburg-Harburg, Thermische Verfahrenstechnik, Eissendorferstrasse 38, 21071 Hamburg, Germany Received 4 September 2001; received in revised form 22 April 2002; accepted 30 April 2002

Abstract The recovery of ethanol from aqueous solutions was studied by using supercritical carbon dioxide. At 333.2 K and 10.0 MPa, vapor /liquid equilibrium data of the mixture CO2/ethanol/water were determined. No azeotrope was observed. Theoretical calculation of equilibrium stages was performed and compared with countercurrent column experiments. Separation of extract and solvent was optimized by multistage solvent distillation. The height of one theoretical stage was found to depend on the ethanol content of the liquid phase. Moreover, flooding point measurements were carried out with ethanol/water mixtures of different composition. # 2002 Elsevier Science B.V. All rights reserved. Keywords: Azeotropic mixtures; Countercurrent extraction; Flooding; Phase equilibria; Stage calculations; HETS

1. Introduction In the field of supercritical fluid extraction (SFE), various researchers proposed the use of supercritical carbon dioxide (CO2) for ethanol recovery and the separation of other volatile organic components from aqueous solutions. Among other applications, literature covers the enrichment of flavor fractions from fruit juices [1] or wine [2], ethanol production [3], separation of impurities from fermentation processes [4,5] and dealcoholization of beverages [6]. Only a few papers cover process evaluation from an engineer-

Corresponding author. Tel.: /49-40-42878-3040; fax: / 49-40-42878-4072 E-mail address: [email protected] (G. Brunner). 1 Present address: BASF AG, D-67056, Ludwigshafen, Germany.

ing point of view. Therefore, it was the objective of this study to investigate the mixture CO2/ ethanol/water, to compare literature data to the present measurements, and to perform thermodynamic stage calculations to demonstrate performance and limits of this technology. Up until the present, vapor /liquid equilibrium (VLE) data of CO2/ethanol/water and its binary mixtures were published over a wide range of temperature and pressure. From an economic point of view, gas extraction of ethanol/water mixtures should yield ethanol of high purity to compete with conventional processes. However, in order to obtain a high solubility of ethanol in the vapor phase, many studies were carried out at conditions of complete miscibility of ethanol and CO2 [7,8]. At these conditions, the phase behavior of the ternary mixture CO2/ethanol/water is of type I [9], as illustrated in the left triangular

0896-8446/02/$ - see front matter # 2002 Elsevier Science B.V. All rights reserved. PII: S 0 8 9 6 - 8 4 4 6 ( 0 2 ) 0 0 0 9 1 - 8

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M. Budich, G. Brunner / J. of Supercritical Fluids 25 (2003) 45 /55

Fig. 1. Schematic phase behavior of the ternary mixture CO2/ethanol/water.

diagram of Fig. 1. With reference to the phase diagram it is easy to understand that anhydrous ethanol cannot be produced at a type I phase behavior. Nevertheless, some researchers concluded that it is impossible to break the azeotrope by using supercritical CO2. Lim et al. [10] assumed that ethanol can be concentrated above azeotropic composition whenever the pressure in the ternary mixture CO2/ ethanol/water is below the critical pressure of the binary mixture CO2/ethanol, equivalent to a phase behavior of type II (right triangular diagram of Fig. 1). One of the first studies reporting the possibility to produce anhydrous ethanol by means of CO2 without adding any entrainer was by Nagahama et al. [11]. Experiments were carried out at conditions of type II phase behavior. Further VLE measurements of the mixture CO2/ethanol/water were performed at Kobe Steel Ltd. in Japan at 10.1 MPa and 313, 323, and 333 K [12]. A solvent-free representation of different VLE data is shown in Fig. 2 in comparison with data at ambient pressure [13]. According to data of Furuta et al. [12] and Horizoe et al. [14], it was not possible to enrich ethanol in the vapor phase above azeotropic composition at 10.1 MPa and either 313 or 323 K, conditions of type I phase behavior. However, no azeotrope was observed at 10.1 MPa and both 333 and 383 K, where the phase behavior of the ternary mixture is of type II. At 10.1 MPa and 333 K, solubility of pure ethanol in CO2 is around 5 wt.% [10]. At 333 K, higher pressures are not recommended due to approach-

ing the point of complete miscibility of CO2/ ethanol around 10.7 MPa. Due to scattering in data at 333 K and 10.0 or 10.1 MPa, respectively, and diverging opinions about the existence of an azeotrope, VLE measurements of the ternary mixture CO2/ethanol/ water were carried out prior to stage calculations and countercurrent column experiments.

2. Experimental set-up and procedure 2.1. Equilibrium apparatus Fig. 3 shows the equilibrium apparatus used in this study that comprised two autoclaves connected in series. The apparatus was designed by Bu¨nz [15] with the objective to enable large vapor phase sampling at conditions of very low vapor phase load.

Fig. 2. McCabe /Thiele diagram of CO2/ethanol/water, solvent-free representation.


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Fig. 3. Phase equilibrium apparatus.

The apparatus was thermostated inside a stirred water bath, covered with insulation balls. Temperature inside the water bath held constant to 9/ 0.2 K. Pressure was adjusted to 9/0.05 MPa. Liquid feed was charged to an evacuated autoclave of 500 cm3 that was separated from a second autoclave (300 cm3) by means of high pressure valves. Carbon dioxide was charged into the apparatus by a compressor, and phase equilibrium was achieved by circulating the vapor phase through both autoclaves for at least 10 h by using an air-operated piston pump (designed and built at Hoffmann-LaRoche, Kaiseraugst, Switzerland). Afterwards, the larger autoclave contained a vapor and a liquid phase in equilibrium, whereas the smaller one contained only the vapor phase. After closing all valves, three liquid samples of approximately 5 g were taken in series from the large autoclave via valve V9. Liquid was separated from gaseous CO2 by means of three cooling traps connected in series. Noncondensable gas was transferred to a burette system by a vacuum pump. The entire content of the vapor phase autoclave (ca. 100 g) was withdrawn via valves V4, V6, and

V7 and solidified in an evacuated flask cooled with liquid nitrogen. The total mass of vapor phase sample was determined by weighing the sampling flask. Sublimation of CO2 was carried out in a freezer at approximately 253 K to obtain a solvent-free vapor-phase condensate at the bottom of the flask. Gas was forced to flow through methanol filled traps to determine the loss of small quantities of ethanol and water from the flask during sublimation of CO2. Experimental results for the liquid phase were reproducible by maximum 9/5% with respect to the CO2 content, and for the vapor phase by maximum 9/10% with respect to the condensed liquid. At least two repeated measurements were carried out to reproduce vapor phase composition. Details on the experimental procedure can be found elsewhere [16]. 2.2. Countercurrent extraction apparatus Countercurrent multistage extraction was carried out in an extraction column of 6 m total height (25 mm ID, equipped with 4 m of Sulzer EX

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packing). Packing elements (specific surface area 1710 m2/m3 [17], voidage 0.86% [17], each element 70 mm high, made of stainless steel 1.4404) were fitted at a 908 turned position in relation to the preceding element. The experimental set-up is shown in Fig. 4. Extraction conditions were set to 333.2 K and 10 MPa. Feed was charged to the middle section of the column by a piston pump. Carbon dioxide entered the column at the bottom. Both feed and CO2 were preheated to extraction conditions before entering the column. Loaded solvent was withdrawn from the top. Solvent and extract were separated by pressure reduction down to 5 MPa. Moreover, solvent distillation was established as proposed by DeFilippi and Vivian [18] and reported by Ikawa et al. [3]. The distillation column used for extract separation had a height of 1.5 m and 35 mm ID and was filled with stainless steel mesh packings (diameter 8 mm). The loaded solvent entered the separator at the bottom. The lower section of the column (ca. 0.2 m) was heated to 303.2 K. By refluxing liquid CO2 to the top of the separator with a reflux ratio of 0.5, the residual ethanol content of the regen-

erated solvent was greatly reduced and ethanol recovery was enhanced. Extract and raffinate were withdrawn continuously from the apparatus. Gas coolers were used to reduce the loss of condensable product with gaseous CO2 at ambient conditions. Extract was partly refluxed to the top of the extraction column. Fresh CO2 was added to replace lost solvent. Samples were repeatedly taken after some hours of constant process conditions. Usually, it took 2 / 3 h to obtain steady state conditions with respect to sample composition after establishing constant mass flow.

Fig. 4. Countercurrent extraction apparatus.

Fig. 5. Flooding point apparatus.

2.3. Flooding point apparatus By means of a flooding point apparatus, maximum liquid and vapor phase cross-section capacity was determined while both phases were in equilibrium. The apparatus was designed by Meyer [17] and is illustrated in Fig. 5. It was equipped with 2 m of Sulzer EX packing with 25 mm ID. Flooding points were determined visually via sapphire windows and by measuring pressure drop. Both phases were charged countercurrently


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to the column by gear pumps and collected in a receiver for the pumps. The entire apparatus was built inside a hot-air cabinet.

3. Materials and analytical procedure The carbon dioxide used in this study had purity higher than 99.95 wt.% (Hydrogas, Bad Ho¨nning, Germany). Ethanol and methanol had a purity higher than 99.8 wt.% (grade Lichrosolv, Merck, Darmstadt, Germany). Deionized water was supplied by the university’s utility station. The water content of a sample was determined by Karl /Fischer titration. Additionally, samples with an ethanol content below 10 wt.% and samples from the methanol filled traps were analyzed by gas chromatography to minimize analytical error. The capillary gas chromatograph ¨ berlingen, Germany) (type 8320, Perkin/Elmer, U with flame-ionization detector was equipped with a polar GC column (Stabilwax, 60 m, 0.25 mm ID, 0.25 mm df by Restek, Bad Soden, Germany) and a guard column (Hydroguard FS, 5 m, 0.25 mm ID; Restek). Helium was used as carrier gas, and the split ratio was set to 1:50 for 0.5 ml of injected sample. All analyses were performed in triplicate.

4. VLE measurements of CO2/ethanol/water Fig. 6 shows VLE data of this study for the ternary system CO2/ethanol/water at 333.2 K and 10 MPa compared with literature data. Experimental results are shown in Table 1. The Ja¨necke diagram [19] was used to illustrate the influence of the solvent-free phase composition on mutual solubility of liquid and solvent. Several papers reported difficulties in reproducing VLE data of this ternary mixture accurately by equations of state [10]. Therefore, empirical relationships were derived to represent the vaporphase (Eq. (1)) and liquid-phase line (Eq. (2)) of the Ja¨necke diagram to ensure reliable stage calculations.

Fig. 6. VLE data of CO2/ethanol/water.

yCO2 100 yCO2 1:5 2 4261:29Yethanol 0:0884Yethanol

(1)

xCO2 100 xCO2

1 2:5 23:6 0:367Xethanol 0:000137Xethanol

(2)

with xCO2, yCO2 are the weight percentage of CO2 in liquid and vapor phase and Xethanol, Yethanol are the weight percentage of ethanol in CO2-free liquid- or vapor-phase sample. Liquid phase data are in good agreement with literature data, except from measurements at high ethanol concentrations. Vapor-phase data by Suzuki et al. [20] show a higher solubility of extract in CO2 at 333.7 K and 10.1 MPa, whereas data at 333.6 K are in good agreement to the present measurements. Furuta et al. [21] reported a lower solubility of extract in CO2. Scattering of literature data is due to different experimental techniques and probably due to difficulties in separating ethanol and gaseous CO2 during sampling. However, good agreement was observed with literature data of the binary mixtures CO2/ethanol [22] and CO2/water [23]. Fig. 7 illustrates the separation factor of ethanol to water as a function of the concentration of ethanol in the solvent-free liquid phase. The separation factor decreased from around 30 at infinite dilution of ethanol in water to approxi-


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Table 1 VLE data of CO2/ethanol/water mixtures at 333.15 K and 10.0 MPa Liquid phase (wt.%) xCO

2

3.821 4.851 4.091 3.583 3.882 4.057 5.914 11.345 16.378 14.921 42.573 43.328 58.868 60.048 60.574

Vapor phase density (kg/m3)

Vapor phase (wt.%)

xEtOH

xH O

yCO

0.000 0.276 0.845 0.852 1.604 7.310 24.403 41.522 49.165 52.244 50.858 50.362 39.925 39.843 39.334

96.179 94.873 95.064 95.565 94.514 88.634 69.684 47.133 34.457 32.834 6.569 6.311 1.207 0.109 0.091

99.757 99.768 99.713 99.723 99.630 99.294 98.424 98.046 97.800 97.906 96.853 96.833 96.083 95.575 95.094

2

2

yEtOH

yH O

0.000 0.018 0.056 0.053 0.110 0.451 1.314 1.695 1.940 1.826 2.957 2.976 3.847 4.415 4.896

0.243 0.214 0.231 0.224 0.260 0.256 0.262 0.260 0.260 0.269 0.190 0.190 0.070 0.009 0.010

2

293.5 294.1 292.3 293.6 297.3 299.7 312.2 313.9 319.4 317.7 332.4 333.2 343.8 350.8 357.7

Data by Furuta et al. [21] are larger, whereas those by Lim et al. [10] are smaller compared with results of this study. Good agreement was found to results of Furuta et al. [12]. By sampling the entire volume of the vaporphase autoclave, vapor-phase density was easily determined during VLE measurements. Vaporphase density is required for hydrodynamic evaluation of a countercurrent flow process. Results are shown in Fig. 8 and Table 1. Eq. (4) represents a simple relationship between load (in g/kg) and vapor-phase density at 333.2 K and 10.0 MPa. Fig. 7. Separation factor of ethanol to water.

mately 1.25 at infinite dilution of water in ethanol. No azeotrope was formed at the conditions investigated. Separation factors are larger compared with data at atmospheric conditions. The line in Fig. 7 was calculated by Eq. (3). aethanol=H2 O

yethanol =xethanol ywater =xwater

0:7328:1 exp(0:0265Xethanol ) (3) with aethanol/H2O is the separation factor of ethanol to water.

Fig. 8. Vapor-phase density of the mixture CO2/ethanol/ water.


rV rCO2 1:3L

(4)

with rCO2 (333.2 K, 10 MPa)/290.2 kg/m3; L is the load in grams of extract per kilogram CO2. Equilibrium data at conditions of low ethanol solubility in CO2 are needed to evaluate feasible conditions for solvent regeneration. VLE data of CO2/ethanol at lower pressures are available in literature [24 /26] and are shown in Fig. 9. The lines illustrate the shapes of the phase envelopes. It becomes obvious that the load of gaseous CO2 at lower pressures cannot be neglected. Ethanol solubility below 0.2 wt.% can only be achieved at pressures close to the vapor pressure of liquid CO2. Unfortunately, large amounts of CO2 will dissolve in the ethanol-rich phase at the same time. Therefore, separation of ethanol and CO2 was carried out by solvent distillation.

5. Calculation of theoretical stages The number of theoretical stages for a predefined separation task were calculated by the Ponchon /Savarit method [27,28]. Details on the method are reported elsewhere [9]. VLE data for the ternary mixture CO2/ethanol/water were represented by the above-mentioned equations. The residual load of CO2 after regeneration was not taken into account by the stage calculation model.

Fig. 9. VLE data of CO2/ethanol at conditions of solvent regeneration.

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Fig. 10 illustrates calculated results that are based on a feed mixture of 10 wt.% ethanol that is separated into an ethanol-rich extract (99.0 wt.% ethanol) and a water-rich raffinate (0.1 wt.% ethanol). The number of theoretical stages was calculated as a function of the extract reflux ratio. Solvent-to-feed ratio increases linearly with increasing reflux ratio. The minimum reflux ratio for this specific separation task was calculated to be 6.2, equal to a minimum solvent-to-feed ratio of around 16, while the minimum number of stages is 10. The minimum number of stages required to achieve a certain extract composition increases with increasing ethanol purity as illustrated in Fig. 11. The required number of theoretical stages decreased rapidly up to a solvent-to-feed ratio around 20, whereas a further increase of the solvent-to-feed ratio above 30 did not reduce the number of theoretical stages significantly. Therefore, in this example working with a feed mixture of 10 wt.% ethanol, ethanol separation should be carried out at solvent-to-feed ratios between 20 and 30. A detailed economic evaluation has to take equipment costs and energy costs for solvent recycling into account. Solvent-to-feed ratios are relatively small compared with other countercurrent gas extraction processes [16]. This is due to large separation factors and a relatively high solubility of pure ethanol in CO2 of around 5 wt.% at the conditions investigated. Instead of the solvent-to-feed ratio, the ratio of solvent flow rate to extract flow rate

Fig. 10. Calculation of the theoretical number of stages.

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to the solvent-to-extract ratio when using a feed with less than 20 wt.% ethanol. Solvent-to-extract ratios remain almost constant at a larger ethanol content of the feed and decrease when the feed composition approaches the desired extract composition.

6. Countercurrent and flooding point experiments 6.1. Ethanol recovery

Fig. 11. Influence of extract quality on process requirements.

should be used as an indicator for operating costs. Since the raffinate is a waste product, the extract is the only product that incorporates all process costs. For the above mentioned separation task, the solvent-to-extract ratio becomes 300 kg/kg at a solvent-to-feed ratio of just 30 kg/kg. Fig. 12 demonstrates the influence of feed composition and the number of theoretical stages on the solvent-to-extract ratio. When the desired ethanol contents of raffinate and extract are set constant, extract flow increases with increasing ethanol content of the feed. Thus, the changing slope of operating lines induces a change in the required reflux ratio for a given number of theoretical stages. The calculation reveals that the production of pure ethanol by means of supercritical CO2 is most expensive with respect

Fig. 12. Influence of feed composition on the solvent-to-extract ratio.

By using feed mixtures of approximately 10, 40, and 94 wt.% ethanol, countercurrent column extraction was carried out to compare experimental with calculated results. During first experiments, conditions in the separator were set to 303.2 K and 5 MPa. For different solvent-to-feed ratios and reflux ratios, the amount of residual ethanol in the raffinate remained always above 4 wt.%. After establishing liquid solvent reflux to the top of the separation column, residual ethanol content of the raffinate phase dropped below 2 wt.%. However, product composition was limited by the limited number of stages achieved in both the separation and extraction column. Zobel [30] proposed to use either an adsorbing material like activated carbon or water to absorb residual ethanol after pressure reduction. Activated carbon was also used by Perrut [6] to remove volatile components from distilled beverages. Other suggestions to improve the separation of volatile components include the use of a trapping substance such as glycerol [31]. This procedure is limited to very few applications when the trapping substance is of further use for the product itself. Distillation of CO2 should be preferred to absorption or adsorption to avoid further regeneration steps or additional waste products. Some researchers did not recirculate the solvent due to the high residual load after separation of extract and solvent by means of pressure reduction [32]. Countercurrent extraction without solvent regeneration is very expensive and will never be established in production scale. Experiments by Ikawa et al. [3] were carried out with a feed mixture of 92.83 wt.% ethanol. Solvent-to-feed ratio was set to 26 kg/kg. Extract


purity above 99.6 wt.% was achieved at a reflux ratio of 12, thus obtaining a raffinate phase with 92.15 wt.% ethanol. This is very close to the composition of the feed mixture itself. Although the study of Ikawa et al. [3] demonstrated the potential of supercritical CO2 to produce pure ethanol, the limit of ethanol recovery and the benefits of solvent distillation with respect to raffinate purity were not fully exploited. 6.2. Evaluation of HETS values When a feed with 94 wt.% ethanol was used, extract with 99.5 wt.% ethanol was produced at a reflux ratio of 4 and a solvent-to-feed ratio of 60 (using 9 kg CO2 per h and 150 g feed per h). According to calculations, 12 equilibrium stages were achieved, equal to a height equivalent to one theoretical stage (HETS) of 0.33 m. Liquid solvent reflux was not required during this experiment because the raffinate was still very rich in ethanol (87 wt.%). During experiments with feed mixtures of low ethanol content, HETS values were found to be in the range of 1 m. HETS values reported by Ikawa et al. [3] were calculated by an empirical model and decreased from 0.75 to 0.48 m with increasing reflux ratio. This was probably due to an improved mass transfer at a higher cross-section capacity. However, it is not likely that the cross-section capacity alone accounts for the change of HETS. Similar HETS values were found by other research groups. Bernad et al. [33] worked with a 1.4 m column of 54 mm ID equipped with Sulzer BX packing. Experiments were carried out at 313 K and 10 MPa with 30 wt.% of ethanol in aqueous solution, used as a continuous phase as well as a dispersed phase. Feed flow was varied at constant solvent flow rates, and HETS decreased from 1.5 m at a solvent-to-feed ratio of 2 /0.5 m at a solvent-to-feed ratio of 10. Extract composition did not exceed 90 wt.% ethanol because the phase behavior was of type I. Lim et al. [34] also worked at 313 K and 10 MPa with a feed mixture of 8.5 wt.% ethanol. HETS values were around 0.45 m when the liquid phase was dispersed and around 0.3 m when CO2 was dispersed in the liquid phase. In contrast to Bernad

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et al. [33], only a very small influence of the solvent-to-feed ratio on HETS was observed at a constant liquid flow rate. The opposite effect was reported by Dondoni et al. [35]. This study was carried out at 313 K and 10 MPa with a column of 1.7 m effective height and 28 mm ID packed with Raschig rings of 4 mm diameter. A feed mixture with 7 wt.% ethanol was charged to the top of the column as a dispersed phase. With increasing the solvent-to-feed ratio from 10 to 25, HETS increased from 0.1 to 0.5 m. Concluding, HETS is influenced by the type of packing, cross-section capacity, the method to evaluate the number of stages, and transport properties of both phases, e.g. viscosity and interfacial tension, that influence mass transfer and backmixing. 6.3. Flooding point measurements When running a countercurrent column at its maximum hydrodynamic capacity, flooding will be observed. Several studies on flooding points of SFE processes involving non-aqueous systems are available in literature [17]. Interesting phenomena can be observed for the CO2/ethanol/water system using a flooding point apparatus as mentioned above. Results are shown in Fig. 13. Maximum cross-section capacity was found to be a function of the ethanol content of the solventfree liquid phase. Changes in flooding behavior at high ethanol concentrations between 100 and 70 wt.% are mainly due to the influence of phase

Fig. 13. Flooding point data for CO2/ethanol/water.

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composition on the density of the phases. With a further decrease in ethanol content of the liquid phase, the influence of viscosity and surface tension of the liquid phase starts to become significant. In fact, flooding occurred almost instantly when pure water circulated in the column and a small vapor phase flow entered the bottom of the column. This happened during experiments carried out with a Sulzer EX packing. A few additional flooding experiments were carried out with a Sulzer CY packing (35 mm diameter) that is characterized by a larger distance of the wire mesh layers. In contrast to measurements with Sulzer EX packing, no flooding occurred at lower cross-section capacities when investigating mixtures of low ethanol content. Therefore, the observed change of flooding behavior was found to be a typical scale-down problem. Furthermore, large HETS values for aqueous solutions may also be due to an unsuitable type of packing that causes backmixing. Further measurements should be carried out using a Sulzer CY packing. No flooding was observed at the extraction tower besides few experiments where the raffinate level went beyond the CO2 inlet. Whenever flooding occurred, the experiment could not be restarted by simply removing some raffinate, but the entire hold-up of the packing had to be withdrawn.

7. Conclusions A reliable technique was developed to determine VLE data of volatile components and supercritical CO2. Furthermore, the need of an improved solvent regeneration was pointed out for continuous SFE processes. Solvent distillation should be established for SFE processes that require raffinate almost free of volatile components. Stage calculation helps to understand possibilities and limits of a process. HETS values were found to be a function of the percentage of water in the liquid phase. Flooding point measurements of aqueous mixtures need to be carried out very carefully due to scale-down problems observed when using close meshed packing material.

The application of supercritical countercurrent extraction to aqueous solutions is limited whenever foaming is observed. Chemical reaction of organic components with CO2 should also be taken into account. Besides the production of pure ethanol, the removal of organic fractions of high market value such as flavor components from fruit juices or of pharmaceutic agents from aqueous plant extracts are promising applications of SFE processes.

Acknowledgements The support of this work by Deutsche Forschungsgemeinschaft (DFG) under grant Br 846/ 15-1 is gratefully acknowledged.

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