Table II.
Run No. 1 2 3 4 5
6 7 8 9
Experimental Data and Calculated Results for Dienediol Bioconversion with Various Feed Schedules and Enzyme Titers Solid Dienediol Residue, Mg./L. Enzyme From 2nd Feed From 3rd Feed Titer, Amount of Feed, Mg./L. DU/Ml. 1st feed 2nd feed 3rd feed Exptl. Calcd. Exptl. Calcd.
18080 1230 910 lOCl0
IOC~O
194 157 147 117 190
970
163
660
191 230 245
loto
950
100 113 153 127 145 100 i19 105 132
d = ((calcd. D - exptl. D)/calcd. D ] X 10070. Mean ualuP of d =
feeds bring about precipitation, and the phase equilibrium determines the composition of the precipitate. T h e dienediol in the solid phase once precipitated does not redissolve (see Figure 1) even though the concentration in solution falls to zero. One can speculate that the phase equilibrium does not apply a t this stage because the dienediol in the solid phase is buried beneath a layer of trienediol. Molecules of dienediol originally on the surface of the crystals quickly leave their position in the lattice and are replaced by trienediol. I t would appear that this is a useful model, which can be used to predict the outcome of incremental feed experiments a t least within a limited range. The mechanism proposed for the behavior of this system seems valid. One value of a valid model is that it can be used to explore wide ranges of certain variables quickly when equivalent experimentation would be tedious. As a n example, in order to study how enzyme potencies affect the completion of this bioconversion a t high substrate levels, the results from a hypothetic feed schedule a t different enzyme titers were calculated on the computer. The feed schedule was: 200 mg. per liter of diene-
03
33.0 27.0 52.0 27.0 70.0 72.0
34.7 32.0 65.0 32.0 61.5 58.0
97
86
n
ii:o
45 36
66.0 0.5'%. Standard derivation of d
61 . O
113.0 132.0
118.0 109.0
n
65.0 57.5 =
45.0
73.57,.
diol dissolved in 2 to 1 solvent fed initially, then four successive feeds of 100 mg. per liter each. One hour was allowed between feeds. I t can be seen from Figure 6 that the residual dienediol is affected very little by enzyme titer. The data given here should not be extrapolated to other systems. T h e solution phenomena for other steroid pairs and even for this pair under other conditions, while alike in principle, are a p t to be quantitatively different. The approach, however, seems valid and worthwhile.
literature Cited
(1) Chen, J. W., Koepsell, H. J., Maxon, W. D., Biotechnol. Bioeng. 4, 63 (1962). (2) Hills, F., M.S. thesis, Bucknell University, 1963. (3) Koepsell, H. J., Biotechnol. Bioeng. 4, 57 (1962). (4) Kutta, W., Z. M a t h . Phys. 46, 435 (1901). RECEIVED for review October 12, 1964 ACCEPTED June 7, 1965 Division of Microbiological Chemistry and Technology, 145th Meeting, ACS, New York, N. Y . ,September 1963.
CATALY'TIC COMBUSTION OF C, TO C, HYD ROCARBONS M .
a.
ACCOMAZZO AND
KEN NOBE
Department of Engineering, University of California, Los Angeles, Calif. HE complete oxidation of hydrocarbons by catalytic means Tis important in both the control of air pollution and the development of fuel cells. As a result, there is considerable activity in hydrocarbon oxidation studies. Recently, a number of investigations on the removal of hydrocarbons a t low concentrations by catalytic oxidation have been reported in the literature (7, 2, 6, 7, 5), 70). T h e literature cited, however, is by no means complete. ' T h e investigation reported in this paper is a study of the effect of chain length, degree of unsaturation, and structure on the catalytic combustion of light hydrocarbons. The hydrocarbons studied were methane, ethane, ethylene, acetylene,
propane, propylene, propadiene, propyne, and cyclopropane. T h e same catalyst bed was used throughout the investigation. Since it was of interest to test the efficiency of the catalyst for a n extended period of time, the initial experimental runs were repeated periodically to determine if there was any decline in the effectiveness of the catalyst. Initial hydrocarbon concentrations and space velocities were varied to approximate conditions expected in automobile catalytic afterburners. Experimental
The experimental apparatus is shown in Figure 1. The reactor was constructed of Vycor tubing 19 mm. in inside VOL. 4
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The catalytic combustion kinetics of methane, ethane, ethylene, acetylene, propane, propylene, propadiene, propyne, and cyclopropane on CuO:A1203 has been investigated at initial hydrocarbon concentrations between 182 and 1450 p.p.m. in the temperature range 140" to 51 0" C., and at gas flow rates of 160, n used to 275, and 525 liters per hour (NTP). Empirical rate expressions of the form r = A e - E ' R T P ~were correlate the experimental data. The correlation considered temperature gradients along the bed and external diffusion of reactant to the catalyst surface. The predicted results agreed within 15% of the experimental results at the higher flow rates between 10 and 80% conversion. Methane was the most difficult hydrocarbon to oxidize and acetylene was the easiest. In general, increase in carbon number decreased the temperature necessary for a given conversion. For a given carbon number, the required temperature decreased with degree of saturation. The combustion products were essentially only carbon dioxide and water. The same catalyst bed was used for approximately 1000 hours with no loss in activity.
TO BROWN ,TEMPERATURE RECORDER
B AIR
GAS
INTAKE
J
INTAKE
Figure 1. 1.
Apparatus 6. Catalyst 7. Thermocouple 8. Furnace
Teflon stopcocks Flowmeter Capillary tube Rotameter Mixing chamber
2. 3.
4. 5.
9. 10. 1 1.
Reactor tube Drying tube 10-ml. buret
8(
-
6C
2 0 t-
d
R 4c 0
I-
6 2c
a w
9
f
3
240
I
260
I
280
MAXIMUM BED
Figure 2.
I
300
320
TEMPERATURE
I
. I
340
350
OC.
Ethylene combustion
Initial concentration of ethylene in p.p.m. 2 1330 a t 5 2 5 lilers/hr. 0 805 a t 525 liters/hr. A 440 a t 5 2 5 liters/hr. 0 8 2 0 a t 275 liters/hr. A 1090 at 160 liters/hr. Calculated d a t a neglecting diffusion Calculated data considering diffusion Points. Experimental d a t a
__-_. ----. 426
I&EC PROCESS DESIGN A N D DEVELOPMENT
diameter. A thermocouple well was placed in the center of the tube. The reactor was packed with ceramic beads 50 cm. deep. Sixteen grams of catalyst (25 ml.) lvere added above these beads. The catalyst bed and approximately 33 cm. of the ceramic bead packing were enclosed in a furnace to maintain the reactor a t various temperature conditions. The temperature distribution in the reactor was measured by iron-constantan thermocouples. One thermocouple was located in the preheater section of the reactor and another a t the entrance to the catalyst bed. Subsequent thermocouples were placed in the bed every 1.25 cm. Temperatures were continuously recorded on a 20-point Brown recording potentiometer. For initial hydrocarbon concentrations of 1000 p.p.m. or less and conversions less than 30%, the temperature gradient across the bed could be maintained lvithin 2' C. At higher conversions and/or concentrations, temperature differences as large as 20" C. resulted. During the reaction the furnace controls were set so that the temperature gradient across the catalyst bed \vas minimized. The air was passed initially through a drying tube containing Drierite (anhydrous CaS04, 8-mesh), 5 cm. in diameter and 34 cm. long. Fresh desiccant was used each day to minimize the water content in the air. The air intake and the hydrocarbon gas were metered by calibrated capillary tubes and a rotameter. The hydrocarbon flow rate was accurately determined by timing a rising soap film in a 10-ml. buret. T h e entering gases were then introduced to a mixing chamber which consisted of several perforated plates aligned to prevent channeling. The ceramic bead packing served as an additional mixing chamber. The hydrocarbon-air mixture was sampled a t the entrance and exit positions of the bed. For each run, carbon monoxide, carbon dioxide, and hydrocarbon concentrations a t the exit position were measured continuously until steady state was reached. Then the bed temperature and the hydrocarbon concentration a t the entrance to the bed were recorded. T h e hydrocarbon concentrations were measured with a Carad flame ionization analyzer and detector. The analyzer output was monitored with a recorder and \\-as reproducible to 1%. The temperature of the hydrogen flame is such that only hydrocarbons are detected. Continuous COZ and CO analyses were made with a Beckman Model 15A infrared analyzer and a MSA Lira 300 infrared analyzer, respectively. In addition to the continuous analyzing equipment, an F and M (Scientific Glassivare Co.) gas chromatograph was used to determine if any hydrocarbon cracking occurred. A 6-foot silica gel column \vas used in conjunction with a thermal conductivity detector for methane, ethane, ethylene, acetylene, and propane. A 3-fOOt column was used for propylene, cyclopropane, propyne, and propadiene. The column was operated at 100" C. and a helium flow rate of 25 ml. per minute. Initially the chromatograph was used for each run. However, since no products other than carbon dioxide \rere found for most of the hydrocarbons and the conversion of hydrocarbon determined by the gas chromatograph agreed with the conversion determined by the hydrocarbon analyzer, the gas chromatograph \vas later used only for spot checks.
280
I
!
300
320
I
I
1
Figure 3.
360
I
I
420
440
1
400 IMAXIMUM B E 0 TEMPERATURE, ' C .
340
380
I
I
I
I
1
Propane combustion
Initial concentration of propane in p.p.m. 0 131 0 a t 525 liters/hr. A 770 at 525 liters/hr. 0 500 a t 525 liters/hr. A 665 c t 275 liters/hr. r 6 5 0 a t 160 liters/hr. Calculated data neglecting diffusion 'Calculated data considering diffusion Points. Experimental data
- - _ -. ----.
,411 hydrocarbons u w e of hlatheson C.P. purity (99.5% minimum) except for propadiene and propyne (95y0minimum purity). T\vo hundred and forty grams of catalyst (SOTocopper oxide and 5OcjC alumina) ivcre prepared in the following way. Four batches of \vet catalyst precipitate (60 grams each) \\-ere made separately and then thoroughly mixed together. The wet precipitate !vas then pressed into the catalyst molds and heated to 1SO' C. in an oven for approximately 8 hours. Each batch \vas prepared in a 3-liter beaker by the same procedure. Baker reagent grade cupric nitrate [Cu(SOa)n.3 H 2 0 , 91.2 grams] \vas added to 500 mi. of distilled water and then mixed \vith 120 grams of \vet Filtrol, grade 90 alumina gel (7SOo HZO). This mixture was heated to 80' C. and stirred continuously. Then a solution containing 42.4 grams of Baker reagent grade KOH in 500 ml. of water was added quickly and the mixture \\-as heated and stirred continuously for 30 minutes. Duiing this period the color of the mixture changed from light !blue to dark blue, to dirty green, and finall>-to dark broivn. Then a solution containing 20 grams of KOH in 500 ml. of Ha0 \vas added and after stirring for S minutes the mixture \vas cooled and the precipitate allowed to sertle. ' l h e precipitate \vas Jvashed five to six times until the filtrate \vas neutral. T h e catalyst l>eller,jwere cylindrical, 0.30 cm. in diameter and 0.25 cm. high. The reactor bed \vas then heated to GOO" C . with air flo\ving through continuously for 48 hours. T h e same catalyst bed was used for all experiments. In this manner? the catalyst >,vastested for approximately 1000 hours. Initial experiments \v.ere periodically repeated and no apparent loss in activity \vas detected. I h e catalyst had a BET surface area of 86 sq. meters per gram and a mean port: radius of 54 A.
I n general, conversion increased with temperature, increased with decrease in flow rate, and increased with decrease in initial hydrocarbon concentration. Methane was the most difficult to oxidize. At temperatures of 400' and 500' C. the conversions a t a floiv rate of 275 liters per hour Lvere 15 and SO%, respectively. Acetylene was least resistant to oxidation, since above 220' C., a t all the flow rates studied, greater than 807, oxidation was achieved. During the experiments with acetylene and propyne, trace amounts of carbon monoxide (less than 50 p.p.m.) were detected. Analysis for aldehydes also shoived trace amounts (less than 20 p.p.m.). X carbon dioxide mass balance, assuming complete combustion, ahvays agreed within experimental error and hence it was concluded that conversion to carbon dioxide !vas essentially complete. Determination of Reaction Rate Parameters. T h e experimental data Tvere correlated by the lvell known design equation for catalytic reactors? Fdx = rdTV, using a rate expression of the form, r = kPITn,where k = A e - E ' R T . Since the temperature distribution for each run was knorvn, the rate parameters could be determined for nonisothermal as well as isothermal conditions with external diffusion effects also considered by an incremental calculation procedure. I n this procedure the reactor was divided into eight sections of equal catalyst mass. For small conversions the design l t h increment may be ivritten equation when applied to t h e j in the form
+
Results and Discussion
Some typical experimental results of the catalytic combustion of methane, ethane, acetylene, propadiene, and propyne are shoLvn in the last column of Table I . Figures 2 to 5 show the experimental data for ethylene, propane, propylene, and cyclopropane combustion plotted as per cent oxidation us. the maximum bed temperature (data are represented by the points), The effect of initial hydrocarbon concentration, maximum temperature of catalyst bed, and flow rate on the extent of hydrocarbon conversion is shown.
where
The rate of reaction is given by rj =
A,-E/nT
j (Pi)"
(3)
At steady state, the rate of hydrocarbon combustion will equal the rate of mass transfer of reactant from the bulk stream VOL. 4
NO. 4
OCTOBER
1965
427
to the catalyst surface. pressed as loo
r
80
-
The mass transfer rate may be ex-
rj = koarn (nj - E'?)
(4)
The value of k, may be calculated by use of recent correlations by Yoshida, Ramaswami, and Hougen ( 7 7 ) . They obtained
t
0 I-
x
I
0 1 1 260 280
300
Figure 5.
I
+ I
I
I
I
400
320 340 360 380 MAXIMUM BED TEMPERATURE, 'C.
I
420
Cyclopropane combustion
Initial concentration of cyclopropane in p.p.m. 0 6 8 0 a t 525 liters/hr. 1 8 2 a t 5 2 5 liters/hr. 665 a t 1 6 0 liters/hr. Calcwlated data neglecting diffusion Calculated data considering diffusion Points. Experimental data
__-_.
_-__ .
Table 1.
Experimental and Calculated Results of Catalytic Combustion of C1 to Ca Hydrocarbons
Q,
Trn, C.
Po X 706, ~ ( l ) ~ , ( l l ) 4, E ) ,
a
Atm.
76
70
70
525 525 275 275 275 275 275 160
429 407 480 441 485 426 485 438
1030
19 12 69 38 73 30 77 55
19 12 67 38 71 30 75 53
20 12 68 39 70 31 73 49
525 525 525 525 525 525 275 160
404 368 372 332 396 337 359 351
1360 1360 740 740 335 335 830 1000
53 22 29 38 12 34 38
52 22 29 8 37 12 34 38
52 22 29 8 39 12 34 33
Acetylene
525 525 525 525 525 525 275 275 160
193 172 192 178 181 151 179 152 184
1090 1090 745 745 395 395 700 700 880
46 19 58 34 61 15 67 19 99
46 19 57 33 60 15 66 19 97
48 16 60 34 62 12 68 17 76
Propadiene
525 525
257 241 216 __. 222 183 198 220 197
680 680 680 250 250 375 865 630
91 65 34 75 24 29 58 61
87 63 34 73 24 29 57 60
78 63 34 70 24 31 50 31
232 202 176 210 193
662 450 286 835 835 835
66 35 19 46 26 16
66 35 19 46 26 16
36 18 48 23 12
Hydrocarbon
Methane
Ethane
L./Hr.
_525 -_
525 525 525 275 160 Propyne
428
525 525 525 275 275 275
180
1030 ....
1450 1450 1090 1090
565 1170
8
68
l&EC PROCESS DESIGN A N D DEVELOPMENT
Since the hydrocarbon concentrations were loiv (less than 0.15%), the values of the viscosity and density of air were used. The diffusivity of the particular hydrocarbon in air was calculated using the method of Hirschfelder, Bird, and Spotz (5). The Lennard-Jones force constants were those given by Flynn and Thodos ( 4 ) . Each of the eight units of the catalyst bed contained 2 grams + 1 - rt; = 2.0, Equation 1 becomes of catalyst. Since
The calculation of the conversion across a n increment of catalyst bed required iteration techniques. The incremental temperatures were determined by placing thermocouples along the catalyst bed. If xo> the initial hydrocarbon partial pressure in bulk stream, F , the inlet molal flow rate of hydrocarbon, T 3 ,the temperature a t j , A , E, and n are known, XS, the final conversion, can be determined. I n the first step, with the given values of no and To, ro is determined by iteration u i t h Equations 3 and 4. \l'hen a value of r l is assumed, Equation 5 is used to determine XI, since xo = 0. Since
71 can be determined and r l calculated. rl, then, is compared with the assumed value of r l . If they agree, the calculation is extended to the next increment. The above calculations were performed on a digital computer. Initially, constants A , E, and n in the empirical rate expressions were approximated by considering isothermal conditions and neglecting external diffusion effects. These values were then used lvith Equations 1 to 6, which consider temperature gradients along the bed and external diffusion to calculate the conversion. By careful examination of the experimental and calculated results, the values of A , E, and n were adjusted and the calculations repeated. The calculated results were again compared with the experimental results. This procedure was repeated until values of A , E, and n were found which provided the best agreement betrveen calculated and experimental results. The values of A , E, and n for each hydrocarbon are given in Table 11.
Table 11.
Hydrocarbon Methane Ethane Ethylene Acetylene Propane Propylene Propadiene Propyne Cyclopropane
Empirical Reaction Rate Parameters
A , Mole/ G.-Sec.-Atm." 5.53 x 102 9.20 x 103 5.82 X 10' 6.67 X l o 2 1 . 2 6 X 10' 6.58 X 10' 4.44 x 100 4.15 X 100 2.12 x 10'
E, Gal.
n
23,000 26,000 18,000 19,000 17,300 17,500 15,000 17,000 16,000
0.9 0.7 0.5 0.2 0.6 0.5 0.3 0.0
0.8
To determine the effect of external diffusion on the combustion reactions, calculations were performed as before, considering temperature gradients along the bed, but neglecting external diffusion. Equation 3, then, becomes
and Equation 4 is omitted from the calculations. Typical experimental data along with calculated data considering external diffusion and temperature gradients along the bed [x(II) ] and calculated results neglecting external diffusion but considering temperature gradients along the bed [.(I)] are given in Table I and Figures 2 to 5. The dashed curves in Figures 2 to 5 represent the calculated results considering temperature gradients along the bed but neglecting external diffusion. The solid curves correspond to the calculated results considering both temperature gradients along the bed and rxternal diffusion. The points are the experimental data. In general, the empirical rate expressions were found to fit the data reasonably \cell a t the higher flow rates of 525 and 275 liters per hour (at 25' C.) a t the concentrations studied. The calculated results usually agreed \+ithin 15yo of experimental results in the range of 10 to 80% conversion a t these flow rates External diffusion reduced the calculated conversion by only 5% a t 8OY0conversion. Below 40% conversion external diffusion could be neglected. T h e results predicted by the empirical rate expressions a t the lo\\est flow rate of 160 liters per hour (at 25' C.) were always higher than the experimental results by as much as a factor of 2. I t was oiiginally thought that this discrepancy was due to a n increase in mass transfer resistance from the bulk stream to the catalyst suiface a t the lower flow rate. However, it has been demonstrated that mass transfer considerations had only a small effect on the conversion. Axial diffusion in the direction of flow was next considered, since a reactant concentration gradient does exist along the bed. Calculations indicated that axial diffusion was negligible. The most probable explanation for the discrepancy between the calculated and experimental results a t 160 liters per hour is nonplug flow. Tht. existence of a radial velocity profile for gases floiting through packed beds has been shown by Schwartz and Smith (8) and Dorweiler and Fahien ( 3 ) . I t was shown that the velocity distribution was essentially independent of flow rate above a n average superficial velocity which corresponded to a modified Reynolds number, R e ' = DpGip(l - e), of about 69. The modified Reynolds numbers corresponding to flow rates of 525, 275, and 160 liters per hour in this investigation were approximately 146, 7 7 , and 45, respectively. Hence, it seems reasonable to assume that the velocity profile for the two higher flow rates was essentially the same. At the lower flow rate. indications n e r e that the profile was altered considerably. Unfortunately, the change in the velocity profile for a modified Reynolds number below 69 has not been given in the literature because of experimental difficulties encountered by the invrstigators. It is of importance to determine the extent to which the values of the empiricaI rate expressions were influenced by the velocity profile. since Flug flow was assumed in the calculations. \Vhen the results of Schwartz and Smith (8) were used. a velocity profile for the bed and flow conditions encountered in this study was assumed. With the values of the rate constant determined by assuming plug flow, the conversion of ethylene was determined considering the existence of a ve-
locity profile. T h e calculations indicated that the correlation of the data a t the higher flow rates of 525 and 275 liters per hour assuming plug flow was valid within experimental error, but was not applicable to the data obtained a t the flow rate of 160 liters per hour. An experimental method often used to determine the conditions under which the rate of catalytic reaction is independent of gas film resistance and mixing phenomena varies the weight of the catalyst bed and reactant feed rate in such a manner as to hold constant the ratio, LV/Q. For the case n # 1,
For constant FV/Q, initial hydrocarbon concentration, and bed temperature, the conversion should remain constant. The size of the catalyst bed (16.0 grams) was held constant for all of the hydrocarbon combustion data. After these experiments were completed the catalyst pellets \cere removed from the bed and then 4.0 grams were placed back in the reactor ( l / d of the original bed). The oxidation of ethane was then studied at a flow rate of 130 liters per hour. Four more grams of catalyst were added to the bed and the flow rate was increased to 260 liters per hour while the temperature, initial concentration, and W / Q were held constant a t their previous values. This procedure was repeated until all of the catalyst was replaced. The results are given in Table 111. The conversion and initial concentration were such that essentially no temperature gradient existed along the bed. For flow rates of 260 liters per hour (at 25' C.) and above, the conversion was constant a t 41%. -4t 130 liters per hour, the conversion was reduced to 36%. These results again show that the rate parameters determined a t the higher flow rates were not applicable to the low flow rate regime. I n the analysis of the experimental data to determine diffusion effects, only external diffusion was considered. At the flow rates of 275 and 525 liters per hour, it seems reasonable to attribute the deviations between the experimental and the calculated results to the neglect of pore diffusion effects in the calculation procedure. Recent experiments in this laboratory with catalyst pellets smaller than those used in this investigation indicated that pore diffusion effects are negligible below combustion temperatures at about 350' C. for C1 to Cs hydrocarbons. A computer program is being developed to take into account the effects of pore diffusion in the catalytic
140
iao
I
I
I
ZM
260
MO
340
1
1
I
1
380
420
460
607
REACTION TEMPERATURE
Figure
6.
I 540
*C.
Comparison of catalytic combustion of C1 to
CI hydrocarbons Flow rate = 5 2 5 liters/hr. 1 . Methane 2. Ethane 3. Ethylene 4. Acetylene
VOL. 4
Initial Concentration = 500 p.p.m. 5. Propane 6. Propylene 7. Propadiene 8. Propyne 9. Cyclopropane
NO. 4
OCTOBER
1965
429
Table 111.
w,
Grams 4.0 8.0 12.0 16.0
Effect of Catalyst Bed Size in Catalytic Combustion of Ethane Q3
Liters/ Hr. 130 260 390 520
W/Q,
G.-Hr./ Liter 0,0308 0.0308 0.0308 0.0308
T,
c.
377 377 377 377
x,
70
Po X 706, ConverAtm. sion 800 36 800 41 800 41 800 41
= mass transfer coefficient, moles per sec. per
n
Pi, Pj+l
= = =
Po
=
$R
s( I )
= = = = = = = = = = = = = = = =
x(1Z)
=
PH
Re
Re' r
combustion of hydrocarbons. T h e results will be presented in a subsequent publication. Effect of Hydrocarbon Structure on Catalytic Conversion. To compare the ease of oxidation of the different hydrocarbons studied, the conversions were calculated for a n isothermal bed using the empirical reaction rate parameters given in Table 11. Figure 6 shoivs conversions calculated for a n initial hydrocarbon concentration of 500 p.p.m. and a flow rate of 525 liters per hour (at 25" C.) as a function of bed temperature. Above 40Yc conversion the curves diverge. T h e upper curve corresponds to the calculated results neglecting external diffusion, whereas the lo\ver curve represents the calculation results when external diffusion is considered. In general, external diffusion reduced the conversion by only 5% a t 80% conversion, and below 40% conversion, diffusion could be neglected. Methane \vas the most difficult hydrocarbon to oxidize; acetylene the least difficult. In general, a n increase in carbon number decreased the temperature for a given conversion. For example, temperatures necessary for 3070 conversion of the saturated compounds, methane, ethane, and propane were 483', 391', and 352' C.: respectively. For a given carbon number, the required temperature decreased with degree of saturation. For example, temperatures necessary for 50% conversion of ethane, ethylene, and acetylene were 391', 284', and 180' C., respectively. These results are consistent u i t h the previous work of Stein (70) and Innes (6). Nomenclature
Arrhenius frequency factor, moles per gram per sec. per a t m . n area of pellet per unit mass, sq. cm. per gram area of pellet per unit volume, cm. -l diffusivity of hydrocarbon in air, sq. cm. per sec. 6 , S, mean particle diameter, cm. activation energy. cal. per mole entering molal hydrocarbon flow rate, moles per sec. mass flow rate per unit cross-sectional area of bed, grams per sq. cm.-sec. molal flow rate per unit cross-sectional area of bed, moles per sq. cm.-sec. reaction rate constant, moles per sec. per gram per atm.n
430
I & E C PROCESS D E S I G N A N D D E V E L O P M E N T
sq. cm. per atm. reaction order partial pressure of hydrocarbon, atm. surface partial pressure of hydrocarbon of the j t h and j l t h increments, respectively partial pressure of hydrocarbon in inlet stream, atm. total pressure, atm. volume f l o ~ vrate, liters per sec. gas constant G/a,$p, Reynolds number of packed beds D,G/!(l - E ) , modified Reynolds number reaction rate, moles per gram per sec. average reaction rate, moles per gram per sec. pellet surface area per pellet volume, cm.-' Schmidt number temperature temperature a t j t h increment Iveight of catalyst bed, grams weight o f j t h and j l t h , respectively, grams conversion experimental per cent conversion per cent conversion neglecting external diffusion effects per cent conversion considering external diffusion effects conversion a t j t h and j 1th increments, respectively
ko
r,,
S
sc
T Tj
LV JVj, T17j+l
>(E)
xi,
xj+
I
=
+
+
+
GREEKLETTERS E
P
4
= catalyst bed porosity = viscosity =
shape factor, 0.91 for cylinders
nj,nj+
= partial pressure of hydrocarbon in bulk stream of
Pp ,
= density, grams per ml. = molal density, moles per ml.
jth a n d j
+ l t h increments, respectively, atm.
literature Cited
(1) Accomazzo, M. A., Nobe, K., Chem. Eng. Progr. Symp. Ser. 5 9 , No. 45, 71 (1963). (2) Anderson. R. B.. Stein. K. C.. Feenan. J. J.. Hofer. L. J . E..
lirschfelder, J. O., 71, 921 (1949). (6) Innes. \V. B.. Duffv. R.. J . Azr Pollution Control Assoc. 11. 369 (7)' Johhson, J. E., Christian, J. G., Carhart, H. FV., Ind. Eng. Chem. 53, 900 (1961). (8) Schwartz, C. E., Smith, J. hl..Ibid.,45, 1209 (1953). (9), Sourirajan, S..Xccomazzo, M. A., Nobe. K., "Actes du Deuxieme Congres International de Catalyse." , . Vol. 11., p. . 2497. Editions Technic! Paris, France: 1961. (10) Stein, K. C., Feenan, J. J., Thompson, G. P., Shultz, J. F., Hofer, L. J . E., Anderson, R. B.: Ind. Eng. Chem. 52, 671 (1960). (11) Yoshida, F., Ramaswami, D., Hougen, 0. A . , A . I. Ch. E . J . 8 , 5 (1962). RECEIVED for review December 28, 1964 ACCEPTED May 24, 1965 Division of Petroleum Chemistry, 149th Meeting, ACS: Detroit, Mich., April 1965. Research supported by funds from the University of California's Air Pollution Research Program.
