250
Ind. Eng. Chem. Process Des. Dev. 1985, 24, 250-255
Robinson, C. G.; Paynter, J. C. Proc. ISEG' 77, SOC.Chem. Ing., London 1971, f 7 , 1416.
RusiAska-Roszak. D. Ph.D. Thesis, Technical University of PoznaA, PoznaA, 1981
Stgpikk-Bioiakiewicz, D.; Szymanowski, J. J. Chem. Techno/. Biotechnol.
Szymanowski, J.; AtamaAczuk, B. H Y & O ~ ~ ~ I I " ~ ~1982, Y 9 , 29. Szymanowski, J.; Blasrczak, J. Chem. Stosow. 1982, 26, 99. Whewell, __.. R. J.; Hughes, M. A.: Hanson. C. J . Inorg. Nucl. Chem. 1975, 37, 2303.
1979, 29. 686.
Stgpniak-Bmiakiewicz, D.; Szymanowski, J. J. Chem. Techno/. Biotechnol. 1981a, 37, 470. Stgpnlak-Biniakiewicz, D.;Szymanowski, J. Hydrometallurgy 198lb, 7, 299. Stqpniak-Biniakiewicz, D.: Szymanowski, J.; Atamahczuk, B. Chem, Stosow.
Received for reuiew April 7, 1983 Accepted April 19, 1984
1983, 27, 249.
Use of Gasoline To Extract Ethanol from Aqueous Solution for Producing Gasohol Fu-Ming Lee' and Robert H. Pshl Phillips Research Center, Phllips Petroleum Company, Bartlesville. Oklahoma 74004
The purpose of this study was to explore the technical feasibility of using gasoline to extract ethanol from aqueous solution for producing gasohol directly. Gasohol is a mixture of 10 vol % ethanol, derived from agricultural products, and 90 vol % unleaded gasoline. Our experimental results from both a simulated &stage countercurrent extraction scheme and a bench-scale continuous extraction column show that gasohol can be produced by extracting ethanol from a 90 wt % ethanol aqueous solution. The extractions were conducted at temperatures from ambient up to 108 O F , and four difterent unleaded gasoline bases were successfully used as the solvents. These results were qualitatively predicted from our theoretical calculations based on the phase diagrams using two different types of gasolines as the extractive sotvents. Various amounts of water will be extracted along with ethanol, depending upon the aromatics content in the gasoline. I n a continuous extraction column, the gasohol extracted by highly aromatic gasoline (33 vol % aromatics) could contain up to 0.7 wt % water. Even a minor amount of water in gasohol will create a phase separation in colder cllmates. Both molecular sieve adsorption and chemical treatment can effectively remove the minor amount of water from the product gasohol. However, the economics of these treatments need to be investigated.
Introduction Gasohol is a mixture of 10 vol % ethanol, derived from agricultural products, and 90 vol % unleaded gasoline. It is one of the more controversial topics of the current energy debate, because some people argue that too much energy is required to process grain into anhydrous ethanol (Chambers et al., 1979; Scheller and Mohr, 1977). For example, the distillation steps alone may consume a little more energy than the fuel value of the product ethanol. Researchers at various institutes are currently working on new technologies to minimize the energy required to separate ethanol from water solutions. One of the novel techniques is to use cellulose or cornstarch to absorb some of the water, so that the energy used in recovering anhydrous ethanol from fermentation broths can be cut to about 10% of the ethanol fuel value (Ladisch and Dyck, 1979; Hartline, 1979). Other new methods include gas chromatography, membrane technology, supercritical extraction using C02 as the solvent, and adsorption using zeolites (molecular sieves). Although there is little doubt that these less conventional separation techniques may prove more energy efficient in the long run, improved conventional distillation, extractive distillation, or liquid-liquid extraction is still the technology of choice for the immediate future. Regarding the liquid-liquid extraction method, we felt that it might be possible to use gasoline to extract ethanol from an ethanol-water mixture. The expectation was that the extract stream would contain 10 vol % or more ethanol and could be used as gasohol. Although some researchers were quite critical of this idea, Leeper and Wanket (1982) o i 9 6 - 4 3 0 5 m i i i i24-0250$01 .5am
at Purdue University recently published their proposed scheme for separation of ethanol from water by extraction with gasoline and presented experimental data on the extraction step. The first part of our investigation was to explore the technical feasibility of the process concept of using gasoline as the selective solvent for ethanol extraction. The second part was to study the process variables of a continuous extractive process for producing gasohol in a bench-scale packed column. In this study, however, no attempt was made to evaluate the economics of this extraction process. Theoretical Calculations As a guide for our experimental work, graphical calculations were carried out on the ethanol-water-gasoline diagrams for a countercurrent multistage extraction process. The phase diagrams of two different gasolines were given by Nakaguchi and Keller (1979). The simplified composition analyses of the gasolines are given in Table I. If we use gasoline F (containing 38 vol % aromatics) at 68 O F , an extract with 8.7 wt % ethanol can theoretically be produced from 80 wt % ethanol-water mixture with higher than 99% ethanol recovery. The solvent to feed ratio was determined to be 8.8. The raffinate contained 98.5 wt % water, 1.5 wt % ethanol, and essentially no gasoline. The number of theoretical stages could not be determined since only a part of the tie (equilibrium) lines were shown in the phase diagram. By use of the available information from the phase diagram, the graphical design procedures were carried out and are presented in Figures 1 and 2. 0 1985 American Chemical Society
Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 2, 1985
20
20
251
1
H
IO
20
30
40
SO
50
GASOLINE,
10
80
90
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Figure 1. The extraction of ethanol from aqueous solution at 68 O F using gasoline F (38 vol % aromatics).
GASOLINE, H T %
Figure 3. The extraction of ethanol from aqueous solution at 68 O F using gasoline E (14 vol % aromatics). 0.08
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F R A C T I O N O F E T H A N O L I N WATER PHASE
Figure 2. Equilibrium stages required for extracting ethanol using gasoline F at 68 OF.
In Figure 1, F represents the feed composition entering the first stage. PAIF, PA2Bl,etc., represent the operating lines, while the tie lines have been omitted for reasons of clarity. To obtain the required theoretical stages for extraction, the values of (F, AI), (Bl, A2),etc., were plotted as the operating line, along with the tie-line data as the equilibrium line in Figure 2. With 80 wt % ethanol in the feed and a 7.0 solvent to feed ratio, four theoretical stages are necessary to produce an extract with 8.7 wt % ethanol. The ethanol recovery was 84%.
Figure 4. Equilibrium stages required for extracting ethanol using gasoline E at 68 OF.
Figures 3 and 4 show the results of using gasoline E (containing 14 vol % aromat,ics) to extract ethanol from the same feed under the same solvent to feed ratio and temperature. Two theoretical stages are required to produce an extract with 7.3 wt % ethanol with 65% ethanol recovery. At a lower ethanol feed concentration (70 wt %) and the same solvent to feed ratio, only 5.5 wt % ethanol can be generated in the extract. Six theoretical stages are required to obtain 59% ethanol recovery. Therefore, the ethanol feed concentration is one of the most important factors affecting extraction. A t lower ethanol feed concentrations, both the ethanol product concentration and
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Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 2, 1985
Table I. The Simplified Composition Analyses (Vol %) of the Extractive Solvents solvent* A B C D olefins 16.3 14.0 36.0 0.0 paraffins 68.4 47.8 32.0 38.0 naphthenes 4.8 5.2 11.3 38.0 aromatics 10.5 33.0 20.7 62.0 total
100.0
100.0
100.0
100.0
E
F
10.8 75.2 75.2 14.0
1.8 60.6 60.6 37.6
100.0
100.0
Solvents A, B, and C are commercial blending stocks for unleaded gasoline. Solvent D is a BTX reformate. Solvents E and F were used by Nakaguchi and Keller (1979).
Table 11. Summary of the Experimental Results for Batch Simulation of a 3-Stage Countercurrent Extraction ethanol concn solv C C D A A
A B
feed,
extr,
wt %
wt%
extr, vol %
water concn in extr, wt %
solv to feed ratio
ethanol recovery in extr, wt %
temp, OF
70 85 70 90 90 92 90
4.7 9.1 7.4 10.3 11.3 12.2 13.3
4.5 8.7 7.1 9.4 10.4 11.2 12.8
0.1 0.5 0.4 0.5 0.6 0.7 0.9
7.0 7.0 7.0 7.0 6.0 6.0 6.0
40 78 76 90 85 93 96
75 75 75 76 72 72 76
recovery are lower even if more equilibrium stages are used. To obtain an extract with 10 wt % ethanol, a feed containing more than 80 wt % ethanol should be used. The other alternatives are the use of extract reflux or higher extraction temperature (68 O F is probably too low). According to Figures 1and 3, gasoline losses in the raffinate can be neglected if the raffinate contains less than 30 and 20 wt % ethanol with gasolines E and F as the solvents, respectively. Other calculations showed that the ethanol recovery increased with increasing solvent to feed ratio or the number of equilibrium stages. Gasoline F is a more effective solvent than gasoline E because of its higher aromatic content. Leeper and Wankat (1982) made extraction calculations based on the equilibrium data using n-heptane, n-hexane, and toluene as the extraction solvents. Their conclusions on the effects of feed concentration, solvent to feed ratio, and solvent type on extraction were similar to ours. They indicated that, with 90 wt % ethanol in feed, most of the desired separation could be done in 5 or less equilibrium stages. Simulation of a 3-Stage Countercurrent Extraction 1. Experimental Details. To verify the results from our theoretical calculations, experiments needed to be conducted with a 2 to 4-stage countercurrent continuous extraction process. We selected a 3-stage countercurrent extraction scheme for experimentation. A schematic diagram of this process is shown in Figure 5. Fresh solvent and feed enter at opposite ends of a series of extraction stages. Extract and raffinate layers pass continuously and countercurrently from stage to stage through the system. The system may be composed of a series of mixers, each with its separate settler, or some form of tray column may be used. In order to simulate a 3-stage countercurrent continuous extraction in the laboratory, an operating scheme proposed by Treybal(1968) may be used. A schematic flow diagram of the proposed scheme is shown in Figure 6 where each circle represents a separatory-funnel equilibrium shake-out, and an ideal stage. The symbols S, F, R, and E are designated for the fresh solvent, feed to be extracted, raffinate and extract, respectively. After sufficient cycles have been carried out so that properties of withdrawn extract and raffinate no longer change, the products E,, E2,EB,R1,R2,
I
hAL EXIRR:
EX-RACT E3
Ell Y,.
?PFF YR,E.> "3
Figure 5. A 3-stage countercurrent extraction process without reflux.
R, t7
i 2
E3
Figure 6. Schematic diagram for the simulation of a 3-stage countercurrent extraction process.
and R3 accurately represent the final results of a continuous scheme corresponding to Figure 5. Three commercial blending stocks for unleaded gasoline and a BTX reformate were used as the solvents to extract ethanol from its aqueous solutions. Simplified analysis for each of the extractive solvents is presented in Table I. For convenience, all the experiments were conducted at room temperature. A standard Karl Fischer method was used to analyze the water content in each stream. Ethanol concentration was measured by a modified gas chromatographic analysis. 2. Results and Discussion. Following the operating scheme in Figure 6, seven experiments were carried out and the experimental conditions and results are presented in Table 11. Initially, solvent C was used as the extractive
Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 2, 1985
Table 111. Cooling Test for Wet Gasohol temp, O F orig final ethanol 75 75 11.3 75 30 8.4 75 0 6.9 75 -16 7.3 75 -90 4.8
compn, w t % gasoline" 88.1 91.3 92.9 92.5 95.1
water 0.60 0.30 0.21 0.21 0.10
253
loss, wt % ethanol water 0.0 0.0 28.4 58.5 37.9 72.6 37.5 70.9 50.5 84.8
"Solvent A.
-
EXTRRCT STRERH ICRSOHOLI
solvent in the simulated 3-stage extraction scheme. The final extract stream (E,) contained 4.7 and 9.1 wt % ethanol when the ethanol contents in the feed stream were, respectively, 70 and 85 wt %. We also found that the ethanol recovery in the extract stream increased from 40 to 78% as the ethanol content in the feed stream increased from 70 to 85 wt % This suggests that the ethanol concentration in the feed stream is a key variable in determining the ethanol content and recovery in the final extract stream. It would be, therefore, desirable to use 90 wt % or higher ethanol as the feed to extraction. A similar prediction was made from our theoretical calculations. For example, with 70 wt % ethanol in the feed and 7.0 solvent to feed ratio, the ethanol content in the experimental fiial extract (4.7 wt %) is similar to the predicted 5.5 w t % with solvent (gasoline)E, despite the fact that the solvents and extraction temperature (75 vs. 68 O F ) are different. It was also predicted that, regardless the type of gasoline used, the feed ethanol concentration should be higher than 80 wt % for producing gasohol. A BTX reformate (solvent D) was also used to extract ethanol from a 70 wt % ethanol solution. The resulting extract stream contained 7.4 w t % ethanol which is substantially higher than that of the extract stream when solvent C was used as the extractive solvent. This is probably due to the high aromatics content in the reformate, which has a strong solvency for extracting ethanol. However, the selectivity of the reformate is no better than that of solvent C because four times as much water was also extracted along with the ethanol (from 0.1 to 0.4 wt
.
VRCUUH J R C K E T E O RNO S I L V E R CORTEO
L-J
R A F F I N R T E STRERH
Figure 7. Bench-scale continuous countercurrent extraction column.
of the ethanol will separate out from gasohol if the wet product is cooled from room temperature to 30 and 0 O F , respectively.
% ).
Solvents A and B are actual unleaded gasoline bases and can be blended directly with anhydrous ethanol for making
gasohol. These two bases were used to extract ethanol from 90 wt % and higher ethanol solutions. At high ethanol feed concentrations such as 90 wt % , solvents A and B can produce an extract stream containing 10 to 13 wt % ethanol with up to %96% recovery. Solvent B showed stronger extraction power than that of solvent A, probably because the former has a higher aromatic content. The discouraging fact is that various amounts of water were extracted simultaneously with ethanol by the extractive solvents. The higher the solvency of the extractive solvent (or the higher the aromatics content in the solvent), the higher the water content in the extract stream. For example, solvent B extracted 0.9 wt % water along with 13.3 wt % ethanol because the solvent contained almost 33% aromatics. This high water content will certainly present phase separation problems in colder climates. To investigate the potential phase separation problems, cooling tests were conducted for a wet gasohol at various low temperatures. The wet gasohol consisted of 11.3 wt % ethanol, 88.1 wt % gasoline (from solvent A), and 0.60 wt % water. The experimental results are summarized in Table 111. The data show that the water content in gasohol has to be less than 0.21 wt % (API Publication 4082 showed 0.17 wt % ) for cold weather application (0 OF). With 0.60 wt % water in gasohol, about 28 and 38 wt %
~
~
Continuous Extraction of Ethanol from Aqueous Solutions in a Bench-Scale Packed Column 1. Experimental Details. The purpose of this part of the investigation was to study the performance of a continuous countercurrent ethanol extraction process. Three different gasoline bases (solvents A, B, and C) were used as the extractive solvents. As shown in Figure 7, a 1.5411. i.d. glass column packed with 3 f t of Goodloe packings was used as the countercurrent extractor. The column was vacuum jacketed and silver coated to minimize heat loss. Both the ethanol feed and the solvent were heated and metered before entering the extraction column. The feed stream for the dispersed phase was fed through a disk-type distributor 1.25 in. in diameter, perforated with 10 0.03-in. holes. Our preliminary study concluded that smaller holes in the distributor would produce very fine droplets that would prevent phase separation in the column. 2. Results and Discussion. The initial runs were made with ethanol feed (heavy stream) as the dispersed phase and the solvent (light stream) as the continuous phase. However, the phase contact was poor and the extraction results were unsatisfactory. A significant improvement in extraction was realized by reversing the phase contact made with the ethanol feed as the continuous phase and the solvent as the dispersed phase. Table IV summarizes the experimental results where the solvent was in the dispersed phase and the ethanol feed was in the continuous phase. From run no. 1 and 2, it
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Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 2, 1985
Table IV. Continuous Countercurrent Extraction of Ethanol from Aqueous ethanol concn feed, extr, extr, water concn solv to run no. solv T, OF wt % wt % vol % in extr, wt % feed ratio 10.9 10.4 0.7 78 88.4 4.3 1 C 83 88.5 9.0 8.3 0.4 5.8 2 A 90.7 11.2 10.3 3 A 99 0.5 5.5 0.5 5.5 90.7 12.1 11.1 4 A 108 12.9 12.4 0.9 3.8 75 88.9 5 B 0.4 3.1 7.6 7.3 6 B 75 83.8 0.2 3.1 76 71.0 3.5 3.4 7 B
Solution for Producing Gasohol” flow rate lb/h feed solv extr 1.03 1.03 1.03 1.03 1.76 1.76 1.76
4.38 5.97 5.65 5.69 6.75 5.46 5.42
raffinate
4.06 5.83
1.07 1.11
6.12 4.80 5.30
2.06 2.07 1.63
ethanol recovery in extr, wt % 48.6 57.6 83.2 91.4 50.6 24.9 15.0
OEquipment 1.5-in. i.d. glass column packed with 3 f t of Goodloe packing. Contact mode: solvent (gasoline) in dispersed phase and ethanol feed in continuous phase.
appears that solvent A would be a better solvent than solvent C, with regard to potential phase separation problems in gasohol, because the former extracts much less water and only slightly less ethanol than the latter. It was also found that the 3-ft column is capable of producing an extract with 10 wt % ethanol as long as the feed ethanol concentration is around 90 wt %. Also, when solvent A was used, ethanol recovery was improved significantly from 57.6 to 91.4 wt % by increasing the extract temperature from 83 to 108 O F (see run no. 2 and 4, Table IV). The water content in extract increased only slightly (from 0.4 to 0.5 wt %). The ethanol recovery can also be improved by using a longer extraction column or higher solvent to feed ratio or both. At higher than room temperature, the light ends of solvent A vaporized and bubbled through the column and disturbed the overhead phase separation. This problem can be eliminated by operating at a pressure above the vapor pressure of the solvent at extractor temperatures. In run no. 5,6, and 7, solvent B was used as the selective solvent for evaluating the extractor performance. The experimental results indicate that, as expected from theoretical calculations, ethanol concentration in the feed stream is a key variable in determining ethanol recovery and ethanol content in the extract stream. Because of the high aromatic content (33%) of the solvent, the extract contained almost 1.0 wt % water along with 12.9 wt% ethanol. As mentioned earlier, this high water content will certainly cause phase separation in cold weather. The ethanol recoveries in the extract stream for these three runs were lower than those of the previous runs because of the lower solvent to ethanol feed ratio (3.1).
Dehydration of Wet Gasohol To eliminate (or minimize) water in the extract stream (gasohol), two different methods for dehydrating wet gasohol have been investigated. The f i t method was to use 3A molecular sieves to selectively remove water from gasohol. The preliminary study was conducted in a l in. i.d. x 12 in. long glass column packed with 1/16-in.3A molecular sieve pellets. A t room temperature, the water content in gasohol was decreased from 0.43 to 0.25 wt % under continuous processing conditions. Ethanol and gasoline components were not adsorbed in the operation. It was anticipated that “dry” gasohol can be produced with a longer adsorption column. Although molecular sieves are very effective in drying small amounts of moisture, they are usually costly. The second method was to use organometallic (OM) compounds to convert water into alcohols. In the preliminary study, aluminum tri-sec-butoxide, [C2H5CH(CH3)O],Al, was added into wet gasohol and it was found that all the water was converted to sec-butyl alcohol as fuel and an aluminum hydroxide precipitate which can be filtered from the gasohol. Other commercially available OM com-
H A T E R FIECYC-E
TO F E R H E N T E Q ETHANOL A YRTER H : X T U R E
?
i U
i
CRSOLlNE.
ETMRNOL
d HATER MIXTURE
Figure 8. A schematic flow diagram for producing gasohol by cooling.
ponenta might also be useful for dehydrating wet gasohol. In using OM compounds for drying gasohol, care must be taken to avoid potential contamination problems. Producing Gasohol by Cooling In searching for alternate methods for producing gasohol, we found that gasohol can be produced by simply mixing 93 wt % ethanol aqueous solution with gasoline at room temperature and then cooling the mixture to lower temperatures. For example, we mixed 28.32 g of 93 wt % ethanol solution with 59.88 g of gasoline base (solvent B) at 74 O F to form a homogeneous mixture. The mixture was cooled to -9 O F for 24 h to allow equilibrium and phase separation. The gasoline phase (gasohol) contained 13.4 wt % ethanol, 86.1 wt % gasoline, and 0.5 wt % water. The yield of ethanol was 27.1% per pass and the ethanol phase (raffmate) can be recycled to recover ethanol and gasoline. We anticipate that, at higher operating temperatures and pressures, a homogeneous solution can be produced by mixing 85 wt % ethanol aqueous solution and gasoline. A schematic process flow diagram is presented in Figure 8. The first advantage of this process is that a simple inline static mixer can be used to replace the expensive liquid extractor (mixer-settler or extraction column). Secondly, the water content in the gasohol will not cause phase separation problems because the product has been cooled and stabilized to the lowest expected temperature where the gasohol is to be marketed. However, cost effectiveness of this process has yet to be determined because of the energy requirements for cooling. Conclusions 1. Theoretical calculations based on two different types of gasolines showed that: (1) higher than 80 wt % ethanol feed concentration is required for producing gasohol; (2) higher than 99% ethanol recovery can be achieved if enough equilibrium stages are provided; and (3) solvent (gasoline) losses in the raffinate can be neglected if the raffinate contains less than 20 wt % ethanol.
Ind. Eng. Chem. Process Des. Dev. 1985, 24, 255-261
2. The experimental results from a simulated 3-stage countercurrent extraction scheme were qualitatively predicted by the theoretical calculations. Gasohol with 90 to 96% ethanol recovery can be produced at room temperature if a 90 wt % ethanol feed is used. 3. Solvency of the extractive solvent is directly proportional to the aromatics content in the solvent. However, the higher the solvency, the higher the water content in the product gasohol. Even a minor amount of water in gasohol will create a phase separation in colder climates. 4. Experiments were also successful for producing gasohol in a 1.5-in. i.d. continuous column. Gasohol with 91.4% ethanol recovery was produced at 108 O F with unleaded gasoline. Increasing extraction temperature from 75 to 108 O F , the water content in gasohol increased only 0.1 wt %. The ethanol recovery can be further improved by use of a longer extraction column. 5. Both molecular sieve absorption and chemical treatment can be effectively used to remove the minor amount of water in the product gasohol. The economics of these treatments need to be investigated. 6. Instead of extraction, contacting and cooling is an alternative method for producing gasohol. Stabilized wet gasohol can be made by simply mixing 85-93'70 ethanol with unleaded gasoline at an appropriate temperature.
255
The homogeneous mixture is then cooled to allow equilibrium and phase separation. Again, the economics of this process need to be studied. Acknowledgment The authors wish to thank G . E. Hays, R. A. Koble (deceased), and Dr. H. S. Chan for some helpful discussions. We also thank Phillips Petroleum Co. for the opportunity to perform this work and for permission to publish the results. Registry No. Ethanol, 64-17-5. Literature Cited Chambers, R. S.; Heredeen, R. A.; Joyce, J. J.; Penner, P. S. Science 1979, 206, 789. Hartline, F. F. Science 1979, 206, 41. Ladisch, M. R.; Dyck K. Science 1979, 205, 898. Leeper, S. A,; Wankat, P. C. Ind. Eng. Chem. Process Des. D e v . 1982, 21, 331. Nakaguchi, G. M.; Keller, J. L. "Ethanol Fuel Modification for Highway Vehicle Use", DOE Final Report, Contract No. EY-764-04-3883, Modific., A003, (DOE ALO, EY-766-04-3683-31) July 1979. Scheller, W. A.; Nohr, 0. J. CHEMECH 1977, 7, 816. Treybal, R. E. "Mass-Transfer Operations"; McGraw-Hill: New York, 1968; p 472.
Received for review May 4, 1983 Revised manuscript received April 16, 1984 Accepted April 27, 1984
Gas-Liquid Mass Transfer Characteristics in a Bubble Column with Suspended Sparingly Soluble Fine Particles Elzo Sada, Hldehiro Kumarawa, Choulho Lee, and Nobuya FuJlwara Department of Chemlcal Englneering, Kyoto University, Kyofo, 606, Japan
To investigate the influence of suspended particles on mass transfer characteristics in a slurry bubble column, physical and chemical absorptions were performed into aqueous slurries of fine calcium hydroxide particles ca. 7 pm in average size. Such mass transfer parameters as volumetric liquid-side mass transfer coefficient k L o a , specific gas-liquid interfacial area a , and hence liquid-side mass transfer coefficient k,' were determined under various electrolyte concentrations, solid concentrations, and gas flow rates, and k Loacould be correlated by the gas flow rate. The volumetric gas-side mass transfer coefficient k@ was determined and correlated by the gas flow rate. The enhancement factors during absorption of dilute carbon dioxide into aqueous calcium hydroxide slurries were compared with the theoretical predictions based on the fllm theory incorporating a finite slurry concept.
Introduction Slurry bubble columns are widely used in industrial practice as absorbers, strippers, and multi-phase reactors because of their simple construction, higher heat and mass transfer coefficients, and good controllability of the liquid residence time. They are typically applied to the processes of coal liquefaction, gas absorption, catalytic reaction, and fermentation. In these processes, the solid particles may be reactants, catalysts, or reaction products. In general, the degree of influence of suspended particles on gas-liquid mass transfer characteristics depends on the particle size, the particle concentration, the liquid-solid density difference, the design parameters of bubble column, the type of gas sparger, and the operating conditions such as gas and liquid flow rates, and others. To date, the influences of the suspended particles on the gas-liquid mass transfer characteristics have been studied by several investigators. Dhanuka and Stepanek (1980) measured the 0196-4305/85l1124-0255$01 .SO10
volumetric mass transfer coefficients in 1.98-, 4.08-, and 5.86-mm glass ballotini suspended bubble columns and concluded that the particles larger than 4 mm increase the volumetric mass transfer coefficients. Kim et al. (1972) reported the existence of the critical particld size below and above which bubbles are likely to coalesce and disintegrate, respectively. That is, they indicated that coalescence occurs when the particles are smaller than about 2.5 mm and disintegration is met when they are larger than 2.5 mm. Kat0 (1963) and Kat0 et al. (1973) investigated the effect of suspended glass beads of 63-177- Mm diameter on the overall volumetric mass transfer coefficient KGa. It was concluded that the KGa decreases with increasing mean particle size and average concentration of solids, but ~ gradually inas the gas velocity increases, K G Cbecomes dependent of both the particle size and the average solid concentration. Sittig (1977) reported a slight increase in the mass transfer coefficient in lower solid concentrations. 0 1985 American Chemical Society
