Kinetics of the Oxidative Dehydrogenation of Butene to Butadiene over a Ferrite Catalyst John S. Sterrett and Howard G . Mcllvried* G u l f Research 8 Developrnenf Company. Pittsburgh. Pennsylvania 15230
A kinetic model has been developed for the oxidative dehydrogenation of butenes to butadiene over a zinc chromium ferrite catalyst. The data on which this model is based were obtained in a plug flow reactor using a set of statistically designed experiments in which reaction temperature was varied from 600 to 680°F, butene mole fraction varied from 0.05 to 0.1, and oxygen mole fraction from 0.031 to 0.062. Total pressure was 1 atm, and several space velocities were run at each experimental condition. The selectivity to butadiene remained high (94-97%) throughout the series of runs even though the experimental work covered a range of butene conversions from 5 to 70%. The rate of formation of carbon dioxide, the only nonselective product found, followed zero-order kinetics. The formation of butadiene, however, did not follow any simple order kinetics and was fit to a semiempirical rate expression using a non-linear, least-squares, computer curve-fitting technique.
Introduction
For many years, butadiene has been manufactured by passing butene over a dehydrogenation catalyst a t appropriate conditions of temperature and pressure (Thomas. 1970). The butene feed is generally diluted with about 12-20 mol of steam per mole of hydrocarbon. both to help stabilize the temperature during the endothermic dehydrogenation reaction and to help shift the equilibrium in the desired direction by lowering the partial pressure of the product hydrogen and butadiene. On-stream periods are generally short. as coke rapidly builds up on the catalyst which must be regenerated every few minutes to every few days, depending upon the catalyst used (Yoge and Morgan, 1972). More recently. oxidative processes have been developed for converting butene to butadiene (Adams. e t ai.. 1964; Ratist. et n / . . 1966: Batist, et ai..1968; Boutry. et ai.. 1969; Cares and Hightower, 1971; Husen, et ai.. 1971; Newman. 1970; Pitzer. 1972; Simons. et a / . . 1971; Wolf, et a / ., 1966). These oxidative processes exhibit several significant advantages over the older processes. chief among these being that when hydrogen is removed from the system by being converted to water, the constraints of thermodynamic equilibrium are moved far to the butadiene side of the equation. For example, the equilibrium conversion for butene to butadiene in conventional dehydrogenation is 35% a t 930°F and 7 1 7 ~at 1110°F at a pressure of 0.1 atm. while the equilibrium conversion in oxidative dehydrogenation is essentially 100% over the complete temperature and pressure range of interest. In practice, nonoxidative commerical processes operate with 4050% conversions and 70-9070 selectivities. The newer oxidative processes can operate at per pass conversions of a t least 7070 with selectivities of 90% or greater. Another advantage of the oxidative process is that much longer on-stream periods are possible than with the nonoxidative system. Oxygen in the vapor phase inhibits the accumulation of carbon on the catalyst so that run lengths are measured in terms of months rather than minutes or hours. Still another advantage of the oxidative system is a more favorable heat balance, for although conventional dehydrogenation is endothermic. oxidative dehydrogenation is exothermic. Because these advantages make oxidative dehydrogenation of butene to butadiene look very attractive. work was performed at this laboratory to develop a n oxydehydro bu54
Ind. Eng. Chem., Process Des. Develop., Vol. 13, No. 1, 1974
tadiene process. Previous papers (Massoth and Scarpiello. 1971; Rennard and Kehl, 1951) have characterized in detail the ferrite catalyst developed and have discussed proposed reaction mechanisms. This paper, which is a n extension of the previous work. presents a kinetic study of the oxidative dehydrogenation reaction. The rate expression which has been developed should prove useful for commercial reactor design. Experimental Section Catalyst a n d Chemicals. The catalyst used in this work was a n unsupported ZnCrFeOd spinel with approximately a 1:1:1 proportion of zinc, chromium. and iron. This catalyst system has been extensively characterized. and since its method of preparation is described elsewhere (Kehl and Rennard. 1969; Kehl, et ai.,1950). this will not be discussed here. Surface area of the catalyst varies from 1 to 10 m2/g depending upon calcination temperature. The kinetic data were obtained on a catalyst with a surface area of' 1.8 m2/g. while the adiabatic run was made on a catalyst with a surface area of 6.2 m2/g. Both catalysts were 10-20 mesh granules and were diluted with 10-20 mesh silicon carbide in a volume ratio of 1: 1. The mixed butene-2 used was supplied by Philips Petroleum Co. and had a purity of 99.0 mol O/C. The oxygen and nitrogen were supplied by Air Products and Chemicals. Inc. All three materials were used without further purification. Equipment a n d Procedure. A schematic diagram of the unit used to generate the kinetic data is shown in Figure 1. Xitrogen and oxygen in a 4:l volume ratio were charged to the unit through calibrated rotometers: water and liquid butene were pumped into the system from calibrated feed tanks. The butene was flashed to atmospheric pressure, mixed with nitrogen in a surge tank, and combined with water and oxygen and fed into a n electrically heated preheater which generated steam and superheated the vapor stream to the desired temperature. The fixedbed, downflow reactor consisted of a 1-in. i.d. stainless steel pipe 32 in. long. The reactor was divided into five zones, each with its own individual electrical heater and air cooler. This provided suitable temperature control so that essentially isothermal conditions mere obtained throughout the catalyst bed. To help avoid hot spots in the reactor, the catalyst was diluted with silicon carbide. The steam and nitrogen in the feed also helped maintain
REACTOR
fU
i/ 1 1
~
0
LlOUlO SEPARATOR
TANK
0
I
I
I
14
TEST METER SAMPLE
I
40 60 EO T I M E ON S T R E A M , h r
20
I
100
I
120
Figure 2. Checkpoint conversions
r-
TO D R A I N
Figure 1. Ychemat ic diagram of experimental equipment
isothermal conditions by acting as heat sinks. Downstream from the reactor, steam was condensed and piped to a drain; the gas product was passed through a wet test meter and vented. On-stream gas chromatographs provided an immediate analysis of the off-gas. Two chromatographs were employed to obtain a complete analysis. In one chromatograph, a column packed with 2-ethoxyethy! sebacate on Chromosorb P separated nitrogen plus oxygen, carbon dioxide, butene isomers, and butadiene. In the other chromatograph. a column packed with molecular sieves partitioned nitrogen and oxygen. Adiabatic runs were made using a 58-in. i.d. reactor. Operating conditions were an inlet temperature of 550"F, a steam-to-butene mole ratio of 20. and an oxygen-to-butene mole retio of 2/3. The reactor contained a 0.25-in.: thin-wall thermowell extending the length of the catalyst bed which permitted determining the temperature profile along the entire catalyst bed.
(kexpr/ho)eespr
C,H,
(1) (2)
Development of Kinetic Model
--
+ 'i.02 + 60. +
=
4C02
Analyses of reactor effluents showed that butadiene, carbon dioxide, and water were the only products formed when butene was oxidatively dehydrogenated over zinc chromium ferrite, catalyst; no carbon monoxide or oxygenated organics were found. The reactions taking place are C.H,
O,,,
where H,, is the space time which would have given the actually observed conversion if the catalyst had been at fresh activity. Since corrections were small in all cases, the fact that finally determined kinetics deviated from first order did not introduce appreciable error. The corrected data from the factorial design are given in Table I. Midway in the experimental program. the catalyst volume was reduced 50% at the same space velocity. thus ince halving the linear velocity through the catalyst bed. 5' this did not significantly change conversion. it was concluded that interparticle diffusional resistance had a negligible effect on the overall kinetics. The effect of intraparticle diffusion was estimated using a method presented by Satterfield (1970). The highest rate of reaction for the range of variables explored in this kinetic study occurred at 680°F at the reactor inlet. Cnder these conditions, the estimated effectiveness factor was 0.95. Since higher effectiveness factors would occur at lower temperatures and at points farther into the catalyst bed, diffusional effects may be neglected in the development of the kinetic model without introducing a significant error.
Results
C,H,
checkpoints at the midpoint of the design were run, and a linear least-squares fit of checkpoint conversions US. run time was made as shown by Figure 2. A rate constant L'S. time on-stream relationship was then developed by assuming first-order kinetics, and a corrected space time was calculated from the following equation
+ H20 + 4H!O K O , + 3H?O
C,H, -----* (3) To determine the extent of thermal reactions and to ensure that the silicon carbide used to dilute the catalyst was truly inert, blank runs were made with the reactor filled with silicon carbide. These runs showed that thermal conversion is insignificant over the temperature range studied, For example, the conversion over silicon carbide at 750°F and a butene GHSV of 450 was less than 1% with no measurable formation of carbon dioxide. The objective of this study was ro develop a kinetic expression which could be used for reactor design purposes. The experimental data were obtained by carrying out a three-level, three-variable, half-factorial design. The variables whose values were changed were temperature. butene concentration, and oxygen concentration. Each experimental condition was run at three to five space velocities; duplicate on-stream gas chromatographic analyses were run to assure t h a t the unit was operating at steady state at each experimental condition. Nitrogen, hydrocarbon, and oxygen balances provided excellent checks on the accuracy of the data. Although coke lay down on the catalyst was quite low, there was some slight catalyst aging throughout the duration of the kinetic runs. To compensate for this. periodic
L_
The rates of formation of butadiene and carbon dioxide are the only expressions which need to be developed, since stoichiometry then fixes the rates of reaction of all the other components in the system. Development of the model was simplified by the discovery that the rate of formation of carbon dioxide follows zero-order kinetics with respect to both hydrocarbon and oxygen over the range of conditions explored in the kinetic runs. This is illustrated by Figure 3 which shows a plot of COZ yield c s . reciprocal space velocity for all the data at 680°F. The four runs shown in the figure correspond to four different concentrations of reactants. Similar zero-order plots represent the experimental d a t a €or carbon dioxide formation at 600 and a t 640°F. When rate constants are determined from these graphs and used in an Arrhenius plot, a very good straight line results with a slope corresponding to an activation energy of approximately 20 kca1,'g mol, as shown by Figure 4. The final value of 20.8 resulted from computer optimization. Because zero-order kinetics fit the data for COz production so well, this form was retained in later computer fitting of the data, and the computer was not given the option of determining the exponent on the partial pressure of hydrocarbon. It is obvious, however. that as the oxygen Ind Eng. Chem., Process Des. Develop., Vol. 13.No. 1. 1974
55
Table I. Butadiene Kinetic Data g
of catalyst
hr g of total
Run
Reaction temu, "F
feed
Butene
1
680
0.000
0,050
0,069 0.144 0,223 0,303 0.398 0.000 0.085 0.172 0.260 0.306 0.000 0,087 0.173 0,248 0.333
0.034 0.030 0.023 0.019 0.015
2
3
4
5
6
7
8
9
10
11
12
Check point 13
680
680
680
640
640
640
640
600
600
600
600
640
0.074 0.148 0.177 0.244 0,000
0.100
0,069 0,138 0.219 0.287
0.091 0.088 0,081 0.077 0.100 0,090 0.085 0.080 0.076 0.050 0.039 0.037 0.033 0.030 0,050 0.042 0.039 0.037 0,035 0,050 0,044 0,042 0.040 0.039 0,100 0,092 0,089 0.087 0.075
0.000
0.056 0.165 0,257 0.346 0,000 0,108 0,194 0.264 0.361 0.000 0.077 0.131 0,204 0,273 0.000 0.128 0.223 0.302 0,392 0.000 0.113 0,256 0.313 0.000 0,091 0.184 0.280 0.000 0 085 0.177 0.246 0.330 0,000 0.237
Steam
047 037 034 028 023 021 047 037 026 023 021 062
0 717 0 727
0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0
0 0 0 0 0
0
0 0 0
0 0 0 0
0 0
0 0
0
0,070
0
0.067 0.066 0.075 0,071 0.067
0 0 0 0 0 0 0
0,066
0.062
051
044 040 033 031 022 018 014 011
031 027 023 020 018 062 056 051
045 045 062 055 052 049 047 031 027 024 022 020 047 043 041 039 040 047 041 039 039 031 028 025 024 062 059 056 056 053
0 047
0.075 0.058
0 035
concentration approaches zero, the rate of reaction must also approach zero. Therefore, a term of the form poz/(l K 0 2 ' p ~ ~was 2 ) introduced to ensure this behavior. A similar term was introduced into the rate expression for the production of butadiene. The formation of butadiene did not follow any simple order kinetics. A number of surface models with adsorption of reactants controlling, surface reaction controlling, and desorption of products controlling were investigated
+
56
Oxveen 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0
0.100
0.086 0.078 0.067 0.064 0.075 0.059 0.050 0,044 0.039 0.075 0.059 0,053 0.049 0.044
0,000
Partial pressures, a t m
--
Ind. Eng. Chem., Process Des. Develop., Vol. 13,No. 1, 1974
0 730 0 735 0 739 0 743 0 667 0 677 0 683 0 692 0 694 0 613 0 627 0 634 0 637 0 642 0 768 0 780 0 784 0 787 0 791 0 743 0 747 0 763 0 754 0 756 0 589 0 596 0 600 0 604 0 607 0 638 0 647 0 649 0 652 0 654 0 793 0 799 0 801 0 803 0 804 0 717 0 720 0 721 0 723 0 724 0 667 0 673 0 676 0 677 0 768 0 773 0 774 0 775 0 613 0 616 0 619 0 620 0 623 0 738 0 743
c02
Butadiene 0.0149 0,0189 0.0247 0,0286 0.0321
0 0 0 0 0
0016
0,0122 0,0207 0,0305 0.0325
0 0 0 0
0021 0031 0048 0055
0,0151
0 0 0 0
0023 0037 0049 0070
0,0233 0.0286 0,0330
0025 0042 0059 0077
0 0015 0 0023 0 0034
0,0146 0,0205 0.0238 0.0292
0 0041
0.0074 0.0111 0.0170 0,0207
0 0 0 0
0014 0020 0025
0,0086 0.0134 0.0184 0.0211
0 0 0 0
0012 0022 0028 0042
0.0105 0,0122 0.0161 0,0190
0 0016 0 0022 0 0030
0.0079 0.0100
0 0007 0 0010
0.0120 0,0146
0 0016 0 0020
0,0056
0 0006
0,0075 0,0091 0.0103
0 0010 0 0017 0 0016
0.0067 0,0099 0,0119
0 0010 0 0015 0 0020
0.0046 0.0071 0.0079
0 0004 0 0007
0.0037 0.0076 0.0077 0,0120
0 0008 0 0011 0 0014
0,0179
0 0023
0
0008
0035
0 0010
0 0021
but did not provide the desired degree of fit. Finally, a semi-empirical, two-site, rate model was tried and found to give the best fit. A total of nine parameters (two activity coefficients, two energies of activation, four adsorption coefficients, and one exponent) in the following butadiene and carbon dioxide expressions were optimized on a digital computer using a pattern search t o minimize the sum of squares of the differences between observed values and calculated values
-
---
68OOF R U N 600'F RUN
0.05L
I
IV
a 0.03 W -I
0
z
0 02
0.01
0.1
0.2
03
04
S P A C E T I M E , hr
0.1
0.2
0.3
04
SPACE T I M E , hr
Figure 3. Carbon dioxide formation c,j. space time ( g of catalyst hrig of' total teed) showing that zero-order kinetics are followed at 680°F: 0 , run 1; 0 , run 2; A , run :3; 0 . run 4.
Figure 5. Comparison of experimental results with kinetic model for typical runs. Points are experimental data and lines are model predictions. Data are for runs 1 and 9, and space time is g of catalyst hr/g of total feed.
Figure 5 shows a typical result and indicates t h a t the model simulates the experimental data very well. The rate expression for the formation of butadiene includes two terms. The first term represents the rate of butadiene formation from butene and the second term provides for the disappearance of butadiene by combustion. As shown by Figure 3, the rate of combustion is not only zero order in hydrocarbon but is independent of the relative amounts of butene and butadiene present. This implies t h a t the sites promoting combustion are always completely covered with hydrocarbon. Since the rate of combustion does not depend upon which species is present, both butene and butadiene must compete about equally for these sites and their rates of combustion must be about equal. Thus, the CO2 resulting from combustion of butadiene will be proportional to the fraction of butadiene in the total hydrocarbon. This accounts for the ratio ~BD/@HD pH) which appears in the combustion term of' the butadiene rate expression. The kinetic study was carried out on a catalyst of 1.8 m2/g surface area. As pointed out under the Experimental Section, surface area can vary depending upon how the catalyst is prepared. With these low surface area catalysts, it was expected t h a t catalyst activity would be directly proportional to surface area, and this was borne out by short runs on several other catalyst charges with differing surface areas. Therefore, within the range of about 1 to 10 m2/g, the above kinetics can be corrected for catalyst surface area by multiplying the rate constant by the surface area of the catalyst being used divided by 1.8.
+
I
155
I
I60 IIT
I65 x IO~;K-~
Figure 1.Arrhenius plot for carbon dioxide formation
I
I 70
Discussion The kinetic model describes the process sufficiently accurately to permit computer simulation of experiments which would otherwise be costly and time consuming to perform. For example, it was possible to simulate adiabatic reactor operation with complete oxygen consumption. Adiabatic operation is particularly important from a commercial point of view because this permits the use of conventional, fixed-bed reactors. If' ext,ensive heat removal is necessary, then tube-type reactors are necessary a t a large increase in cost. It should also be noted that the extent of the exothermicity is highly dependent on selectivity because of the difference in heat of reaction for dehydrogenation and combustion. For example, while the heat Ind. Eng. Chem., Process Des. Develop., Vol. 13, No. 1 . 1974
57
~~
5 wfn'
-a W
a 3
~
400 -
300
-
a a P W
3 u
200
I-
2 -0a
100
a
RATE OF C o p FORMATION
20 40 60 80 100 % BUTENE CONVERSION AT COMPLETE 02 CONSUMPTION
Figure 6 . Predicted temperature rise for adiabatic operation. Initial butene mole fraction is 0.043. The only other components in
the system are oxygen and steam.
d -----I
'Oo0
ed a severe test for the model. In order to check its adequacy, a run was made in bench-scale equipment a t as close to adiabatic conditions as possible. The results from this run are shown in Figure 5 which illustrates that the profile from the experimental run agrees quite well with that from the computer simulation. The outlet temperature of the experimental reactor was lower than the outlet temperature from the computer simulation because of unavoidable heat losses with a small diameter reactor. In comparing the simulation data with experimental results, the computer simulation predicted 96% selectivity to butadiene whereas experimental results showed 92% selectivity to butadiene. This difference between predicted and experimental results, however, is not large considering the high temperatures (>900"F) achieved in the adiabatic operation compared t o the isothermal temperature levels (600-680°F) from which the kinetic model was derived, and the large difference in surface area, 6.2 us. 1.8 m2/g, between the catalyst used for the kinetic work and that used for the adiabatic runs. Conciusion Oxidative dehydrogenation of butene over a zinc chromium ferrite catalyst appears to be an attractive and flexible continuous process for producing butadiene; conversion is higher than that obtainable with conventional dehydrogenation and selectivity is equal or better. A rate model has been developed which very satisfactorily represents the kinetic behavior of the system and which can be used for reactor design studies.
900
800
600
Acknowledgment The authors wish to thank the B. F. Goodrich Company, which provided part of the financial support for this work. for permission to publish this paper.
500
Nomenclature 0.0
0.2 0.4 0.6 0.8 1.0 FRACTIONAL LENGTH OF CATALYST B E D
Figure 7. Reactor temperature profiles for an adiabatic operation with an oxygen-to-butene mole ratio of 213 and a steam-to-butene mole ratio of 20.
released by the oxidative removal of hydrogen from butene to form butadiene is about 60 kcal/g mol of oxygen, the heat released by complete combustion to carbon dioxide and water is considerably higher 100 kcal/g mol of oxygen. Furthermore, if selectivity is lower, additional oxygen will be required to achieve the desired per pass conversion. thus further increasing the heat load. High selectivity is, therefore, essential not only for good yields, but also for good temperature control. A series of computer simulations were made for adiabatic runs a t four oxygen-to-butene feed ratios for a n initial butene mole fraction of 0.043 with the only other components in the system being oxygen and steam. The maximum temperature rises from these results are plotted in Figure 6 us butene conversion. To ensure that a slight error in the rate of carbon dioxide formation would not lead to a drastic change in selectivity under adiabatic conditions or result in an unacceptably high temperature rise, a second series of simulations was run in which the rate constant for the formation of carbon dioxide was doubled. The temperature rise for these runs is also shown in Figure 6. In both cases, temperature gradients are moderate, even for 60-70% butene conversion. Adiabatic computer simulations represent a gross extrapolation from the conditions at which the kinetic data were obtained, and these adiabatic simulations represent58
Ind. Eng. Chem., Process Des. Develop., Vol. 13, No. 1 , 1974
AI,A2 = frequency factor E1,E2 = activation energy. cal/g mol k = rate constant, hr-I K,K' = adsorption coefficient, a t m - I p = partial pressure, atm r = reaction rate, g mol/hr/g of catalyst R = gas constant, cal/g mol "R T = temperature, O R 0 = space time, g of catalyst hr/g of total feed
Subscripts B = butene BD = butadiene = oxygen cor = corrected expt = experimental 0 = initial 0 2
Literature Cited Adarns, C. R., Voge. H. H., Morgan, C . Z . , Arrnstrong, W. E., J. Catal., 3. 379 119641. Batist. P. A..Lippens, B. C., Schuit, G. C. A,, J. Catal.. 5 , 55 (1966). Batist, P. A., Der Kinderen, A. H . W. M . . Leeuwenburg, Y . , Metz, F. A . M . G.. Schuit, G .C. A , , J . Catal.. 12, 45 (1968). Boutry, P.. Montarnal, R., Wrzyszcz, J . , J. Catai., 13, 75 (1969). Cares, W. R.. Hightower, J. W., J. Catal.. 23, 193 (1971) Husen, P. C . . Deel, K . R., Peters, W. D., Oil Gas J . , 69, No. 31, 60 (19711. Kehl, W . L., Rennard. R . J.. U. S. Patent 3,450,788 (1969). Kehl, W. L.. Rennard, R. J.. Swift, H. E.. U. S. Patent 3,527,834 (1970). Massoth, F. E., Scarpiello, D. A,. J . Catal.. 21, 294 (1971). Newman, F. C., Ind. Eng. Chem.. 62. No. 5, 42 (1970). Pitzer, E. W . , lnd. Eng. Chem.. Prod. Res. Develop., 11, 299 (1972) Rennard, R. J., Kehl, W . L., J . Catai.. 21. 282 (1971). Satterfield, C. N . , "Mass Transfer in Heterogeneous Catalysis," MiT Press, Cambridge, Mass., 1970, Chapter 4. Simons, T. G. J.. Houtman, P. N.,Schuit, G . C. A., J . Catal.. 23, 1 (1971)
Thomas, C. L.. "Catalytic Processes and Proven Catalysts." Academic Press, New York, N. Y., 1970, p 46. Voge, H. H., Morgan, C. 2.. lnd. Eng. Chem.. Process Des. Develop.. 1 1 , 454 (1972). Wolf, C. N.. Bergman, R. I., Sittig, M., Chem. Week, 98, No. 22, 113 (1966).
T h i s p a p e r was p r e s e n t e d a t t h e 4 t h J o i n t AIChE-CSChE C h e m i c a l E n g i n e e r i n g Conference. V a n c o u v e r , C a n a d a . S e p t 1973.
Received for rvripu. May 16, 1973 A c c e p t c d A u g u s t 20, 1972
An Electrochemical Device for Carbon Dioxide Concentration. I. System Design and Performance Jack Winnick,"' National Aeronautics and Space Administration, Johnson Space Center Houston, Texas
Richard D. Marshall, and Franz H. Schubert Life Systems, l n c
Cleveland, Ohio
A system comprised mainly of 90 electrochemical cells has been designed for use a s a carbon dioxide (CO2) concentrator in a manned spacecraft. Cabin gas, with a COz partial pressure of about 3 r n m , I S
passed across t h e air cathode of an electrochemical cell. It is concentrated through the carbonate electrolyte and expelled into the hydrogen-filled anode cavity. The total system, a s well a s the individual cell design, is described. Experimental results are shown for t h e full (90-cell) system and also for smaller scale ( 1 - and 3-cell) tests..Excellent consistency among the tests was found.
Introduction Manned missions in space for extended periods require recycling of most consumables. The recovery of water ( H 2 0 ) and oxygen ( 0 2 ) becomes attractive for missions of about 30 man-days or longer. Oxygen can be regenerated from the metabolically produced carbon dioxide (COz) by three main paths: using direct electrolysis of COz in a solid electrolyte 0 2 regeneration subsystem forming 0 2 and solid carbon; using the Bosch reaction, converting C 0 2 and hydrogen (Hz) through a series of steps to 0 2 , solid carbon, and H20; and using the Sabatier reaction converting COz and Hz to methane and HzO. The product H 2 0 formed by the Bosch and Sabatier reactions is electrolyzed to form make-up 0 2 and Hz. Both the Bosch and Sabatier reactions require H2, although in different ratios to the CO2. Carbon dioxide concentrating systems developed for current ( e . g . , Skylab) manned spacecraft are cyclic adsorption beds utilizing commercial zeolites (Dell'Osso, et al., 1969) as the CO2 sorbent. Criteria for cabin COz environments for future missions require control a t or below CO2 partial pressures (pCO2) of 3 mm. At these low partial pressures the zeolite systems are inefficient resulting in a high system equivalent weight. An electrochemical C 0 2 removal device developed for aircraft (Wynveen and Quattrone, 1971) proved applicable for spacecraft (Wynveen, et al., 1972) and is presently under development. The electrochemical concept offers these advantages: (1) continuous operation instead of cyclic operation; (2) low pCOz capability with low system equivalent weight; (3) concentrated COz, noncontaminated by cabin air, thus protecting the COz reduction process catalyst (Sabatier) or preventing buildup of nonreacting
' Permanent Address,
Department of Chemical Engineering, University
of Missouri, Columbia, Mo. 65201.
gases in C O z reduction recycle loops (Bosch and solid electrolyte); (4) supply of COz premixed at proper ratios with Hz for the Bosch or Sabatier process; ( 5 ) elimination of COz compressors since vacuum desorption is not required. Extensive development experimentation was performed to identify performance parameters using single electrochemical cells, a three-cell unit, and a six-man capacity, 90-cell system. The test results provided a basis for system sizing and performance prediction. The data also served to provide a basis for analytical simulation of' the CO2 removal process (Lin and Winnick, 1974).
Process Description T h e electrochemical reactions occur on electrodes separated by a n aqueous carbonate solution supported in a thin asbestos matrix. The cathode and anode are comprised of fine mesh screens upon which a Teflon and platinum mixture has been applied. Moist air containing COZ at ambient concentrations is passed over the cathode. Oxygen in the air diffuses into the liquid film a t the electrode and reacts with HzO t o form hydroxyl ions (OH-) O? 2H,O de$OH(1)
-
+
+
The COz also diffuses into the liquid, reacting with OH- to form carbonate ions ( C o s 2 - )
CO, HC0,-
+
+
OH-
OH-
-
HC0,-
CO,,'-
+
('a)
HLO
('11)
Moist H2 is passed over the anode. After diffusing through a thin film of solution, it reacts a t the electrode to form hydronium ions (H30+ ) H?
+
~ H ~ -+ O
+
?e-
(3)
The lowered p H in the anode region favors higher bicarbonate ion concentrations via the equilibrium reaction Ind. Eng. Chem., Process Des. Develop., Vol. 13, No. 1, 1974
59
