Ethylbenzene Dehydrogenation Kinetics in a Fixed Bed and a

6 Ethylbenzene Dehydrogenation Kinetics in a Fixed Bed and a Continuous Stirred Gas-Solid

Downloaded by PURDUE UNIV on May 29, 2016 | http://pubs.acs.org Publication Date: June 1, 1975 | doi: 10.1021/ba-1974-0133.ch006

Reactor H. W. G. HEYNEN and H. S. VAN DER BAAN Department of Chemical Technology, University of Technology, Eindhoven, The Netherlands

For the dehydrogenation of ethylbenzene to styrene, aluminasupported uranium dioxide is a very active catalyst with a selectivity above 95% at high conversion. The kinetics of this dehydrogenation reaction can be represented by the following equation: r = k(c-c k )/1+K c+K c ScH/

eq

E

s

s

for temperatures from 465° to 495°C. The kinetic constants in this equation have been derived from experiments in a reactor that could be operated either as a fixed bed plugflowreactor or as a continuous stirred gas-solid reactor (CSGSR). When taking into account the small deviation from ideality of the latter reactor the kinetic constants obtained could describe the results in both reactors within the experimental error.

T

he most important process for styrene manufacture is the dehydrogenation of ethylbenzene, usually carried out at ca. 6 0 0 ° C over metal oxide catalysts in a steam atmosphere. During dealkylation experiments of toluene (1, 2) and ethylbenzene uranium dioxide was an active and stable dehydrogenation catalyst (3), especially when used on an alumina support. Compared with data on commercial catalysts (4, 5) we found that the supported uranium dioxide catalyst showed a markedly improved selectivity even at high conversion levels. This is caused by the high activity of the catalyst which allows reaction temperatures as low as 4 5 0 ° C. In view of these attractive properties and because only little kinetic data on the catalytic dehydrogenation have been published (6, 7, 8), we decided to make a careful kinetic study of this reaction. We started our study of the ethylbenzene dehydrogenation reaction over supported uranium dioxide in a plug flow fixed bed reactor. Experiments with a plug flow reactor however can be inaccurate because mass and heat transport effects may go unnoticed. Very high feed flow rates will alleviate these problems, but the resulting low conversion will introduce inaccuracies in the analysis. We therefore decided to build a continuous stirred gas-solid reactor (9) 67

Hulburt; Chemical Reaction Engineering—II Advances in Chemistry; American Chemical Society: Washington, DC, 1974.


68

CHEMICAL REACTION ENGINEERING

Π

Figure 1. Apparatus (CSGSR) where high internal flow rates could be combined with high con version of the outside feed. We were confident that the combined information from plug flow and CSGSR experiments, as advocated by Kiperman (10), would allow us to determine the kinetic constants more accurately. Moreover, search programs for kinetic constants from CSGSR experiments are less com puter time consuming than those for plug flow experiments. In our CSGSR we used forced circulation of the gas phase through a stationary bed because this allows easy measurement of the bed temperature and of the pressure drop over the bed. The latter information is especially important because from these data the flow rate through the bed can be calculated. We can thus determine how much the behavior of our CSGSR deviates from that of an ideally mixed reactor. For the forced gas circulation we used a centrifugal fan, driven by a magnetically coupled variable speed drive according to Brisk (II). Experiments with our first CSGSR, reactor A, showed that there were discrepancies between the performance of this reactor and that of the plug flow reactor which could not be explained by the differences theoretically expected for both reactors. These discrepancies were most probably caused by differ ences in the history of the catalyst in the two reactors. To ascertain that the conditions in the plug flow fixed bed and the CSGSR were fully comparable we developed a new reactor (reactor B) in which reac tions under both plug flow and CSGSR conditions could be carried out on the same quantity of catalyst and that could be switched over easily from the one mode of operation to the other. Catalyst holder A could not be used for plug flow experiments because the outside gas feed causes a very low linear gas velocity in the bed. Therefore under the applied reaction conditions almost complete backmixing will occur (D /uL — 1). a

Experimental Preparation and Properties of the Catalyst. Uranyl acetate (20 grams ( C H C O O ) U 0 2 * 2 H 0 , Merck, analytical grade) is dissolved in warm water. An excess of concentrated ammonia solution is added, and the yellow precipitate U 0 · x N H · t/H 0 is filtered off and washed with water. Aluminum nitrate, 409 grams ( A 1 ( N 0 ) · 9 H 0 , Merck analytical grade) is dissolved in water and poured into 300-ml concentrated ammonia solution. The precipitate is fil3

3

2

3

2

2

3

3

2


6.

H E Y N E N A N D V A N DER B A A N

Ethylbenzene Dehydrogenation Kinetics 69

tered off and washed with water. The two precipitates are transferred to a flask, mixed with 1 liter of water and kept at 9 5 ° - 1 0 0 ° C under vigorous stirring. After 20 hrs of stirring the solid is filtered off, washed with water and dried for 24 hrs at 135°C. Finally the product is calcined at 6 0 0 ° C in air for 24 hours, and the resulting catalyst is broken and sieved. The specific B E T surface area, using nitrogen adsorption, was 144 m /gram, and the average pore diameter was 46A. Freshly prepared catalyst is reduced to UO2.0/AI2O3, the active dehy drogenation catalyst, by ethylbenzene under reaction conditions. Apparatus. Figure 1 shows a diagram of the apparatus. A constant flow of carrier gas (nitrogen or carbon dioxide) passes through a heated vaporizer V, filled with ethylbenzene or a mixture of ethylbenzene and styrene. The gases are carefully freed from oxygen by BASF R3-11 catalyst. The carrier gas-hydrocarbon mixture passes either via preheater P H through reactor R to sample valve Sj (12) or flows directly via bypass Β to that sampling valve. The reactor feed or the reaction products are introduced in an analysis system consisting of two gas chromatographs, G L C and G L C separated by cold traps. On G L C benzene, toluene, ethylbenzene, and styrene can be deter mined; on G L C , hydrogen is measured. Figure 2 shows the stainless steel


2

X

2

X

2


70


II


CSGSR including the stirrer drive unit. In reactor A, which was used in the preliminary experiments, the annular catalyst holder A (Figure 3, left) was used while the final kinetic measurements were carried out in reactor Β with catalyst holder Β (Figure 3, right).

20

Figure 3.

Catalyst holder A (left); catalyst holder Β (right)

ΔΡ

5000

10000

— · - stirrer speed rpm

Figure 4. Internal recirculation number and pressure drop over the catalyst bed vs. stirrer speed at 500°C. Catalyst holder Β filled with 8.0 grams catalyst (0.6-1 mm); cross sectional area catalyst holder, 3.1 cm*; carrier gas, carbon dioxide, 4.8 liters/hr NTP. Results and Discussion Reactor Characteristics. The degree of mixing in stirred gas solid reactors is usually assessed by tracer techniques. From step response measurements we concluded that reactor A can be operated as a perfect mixer under our experimental conditions. However, mixing in the catalyst bed of reactor Β could not be judged accurately by the step response method because in this case the volume of the catalyst bed (8.5 cm ) is much smaller than the total free volume of the reactor (300 cm ). 3

3


6.

HEYNEN AND VAN DER BAAN


Another way to check the mixing behavior is to measure the pressure drop over the catalyst bed and to calculate the gas flow through the bed by the Ergun relation. With this relation first the actual porosity of the catalyst bed was calculated from measurements of the pressure drop as a function of the feed rate through the fixed bed (central plug closed). The number of internal recirculations η is defined as:


^ _ gasflowthrough the bed _ ^ reactor gas feed

(1)

In Figure 4, for reactor Β, η and the pressure drop over the catalyst bed ( = discharge pressure of the fan) are given as a function of the stirrer speed. The kinetic experiments were carried out in reactor B, which behaved as a good plug flow reactor at the reactor feed rates used during fixed bed experi ments (D /uL = 0.03-0.008). By using a heavier gas (carbon dioxide instead a

I Ι

Ο.

50.

40

10000

5000 1

2 3

10

rpm — -

15 18

Figure 5. Styrene productivity vs. stirrer speed. Reaction conditions: reactor Β filled with 8.0 grams catalyst; reaction temperature 480°C; car rier gas, 5.0 liters/hr NTP carbon dioxide; ethyl benzene feed gas concentration, c = 0.159 mmolej liter. 0

of nitrogen), enough recirculations were provided at high feed rates to approach an ideally mixed reactor. Conversion experiments with varying stirrer speed confirmed this. Figure 5 shows that styrene production as a function of the stirrer speed passes through a maximum and tends to remain constant above 6000 rpm for a reactor gas feed of 5 liter/hr N T P . For the highest feed rate, 8000 rpm is required to approach this condition. At very low η the styrene production is low because bypassing through the central hole occurs while at intermediate η the reactor behaves as a plug flow reactor with a low number of recirculations. Reaction Kinetics. Calculations according to Hougen (23) and Satterfield and Sherwood (14) have established that under the experimental conditions no limitations are to be expected by bulk gas phase or pore diffusion. This was also proved by the results of experiments which showed that a fourfold increase of the average catalyst diameter did not affect the conversion (see



72


500

1000

II

W/F gcat.sec/l

Figure 6. Ethylbenzene conversion vs. W / F at var ious catalyst diameters d : • 0.6 < d < 1.0 mm Ο 0.3 < d < 0.42 mm φ 0.15 < dp < 0.30 mm Reaction conditions:fixedbed reactorfilledwith 2.0 grams catalyst; carrier gas, nitrogen; ethylbenzene vapor pressure in feed gas, 7.3 mm Hg. p

P

P

1000

2000

W/F gcat. sec/I

Figure 7. Ethylbenzene conversion vs. W / F for var ious amounts of catalyst: • 4.0 grams catalyst φ 2.0 grams catalyst Ο 2.0 grams catalyst diluted with silicon carbide Reaction conditions:fixedbed reactor; carrier gas, nitrogen; ethylbenzene pressure in feed gas, 7.3 mm Hg. Figure 6). Figure 7 shows that the conversion of ethylbenzene only depends on the time of contact with the catalyst—i.e., equal W/F values give the same conversions when the amount of catalyst is varied. From these data follows that at constant W/F a variation of the amount of catalyst or dilution of the catalyst



6.


HEYNEN AND VAN DER BAAN

with silicon carbide has no effect on the conversion. Moreover, the conversion of ethylbenzene in the empty reactor and in the reactor filled with silicon carbide (which was in some experiments used to dilute the catalyst) was shown to be negligible under our experimental conditions. A 160-hr experiment under conditions similar to the conditions of the kinetic experiments showed that no decrease in catalyst activity occurs. Therefore, either the chemical reaction at the catalyst surface or adsorption of ethylbenzene or desorption of products is rate controlling. We visualize that in the reaction step an adsorbed ethyl benzene molecule loses hydrogen, leaving an adsorbed styrene molecule on the surface, which will desorb subsequently. As far as the hydrogen is con cerned two pathways are open: (a) The hydrogen is liberated as such to the gas phase (b) The hydrogen is first adsorbed at an adjacent site and subsequently released to the gas phase. If the chemical reaction is rate determining, the two pathways lead to the following rate equations: _

dc _ k(c — CsCn/keq) dW/F ~ 1 + ΣΚια dc

dW/F "

k(c — CgCu/fceq)

(1 +

ZKid)*

(2) (3)

Rate expressions similar to Equation 2 were also proposed by Balandin (15) and Carra (16). Because of the high reaction temperatures the adsorption constants of hydrogen and carrier gas should be considerably lower than the corresponding constants for hydrocarbons and are therefore omitted in the rate equations. At an early stage of this study we found from differential measurements that of the two, only Equation 2 could describe the data.


74


II

When for differential experiments 1/r is plotted vs. l/c , straight lines are obtained if Equation 2 is valid; if Equation 3 is valid, a linear relation exists between ( c / r ) and c . We carried out differential experiments in a plug flow reactor with a small amount of catalyst. The results, plotted in Figure 8, show that Equation 2 can describe the data. These experiments, however, were done with catalyst from a batch that has not been used in further studies. Therefore the kinetic constants that can be obtained from Figure 8 have not been taken into account in this study although they are comparable with those obtained later on. Since no appreciable amounts of byproducts were found (selectivities for styrene varied from 95-99% ), Equation 2 can be changed into: 0

G

1/2

0


ι

des __ k(c — cs — cscn/keo) dW/F 1 + Kv(c - c ) + K c* 0

0

s

/ \ ' 2a

K

8

The equilibrium constant k was calculated from data given by Boundy (17). For experiments with nitrogen as a carrier gas the amounts of styrene and hydrogen detected in the ethylbenzene dehydrogenation products are equal. However, as explained before, carbon dioxide was used as carrier gas in the latter experiments, and thus the amounts of styrene and hydrogen are no longer equal since hydrogen reacts with carbon dioxide. This means that to integrate Equation 2a an additional expression for c is required. We found that for both the fixed bed and the recirculation reactor at each reaction temperature the n/ o t i ° could be described by an empirical relation of the form: eq

H

c

c

r a

Table I.

Reaction Conditions for Kinetic Experiments in Fixed Bed and in C S G S R β

Vapor Pressure in Feed Gas, mm H

g

I

Reaction Temperature, °C 465 480 495

Ethylbenzene 3.0 7.5 8.55 14.0 30.3

Styrene 0 0 6.1 0 0

CCI* KJILO L VULV, liters/hr NTP 5 7.5 10 20

* Feed gas: carbon dioxide, purity 99.995%; catalyst: 8.00 grams, 20 wt % U O 2 on AI2O3 0.6—1 mm; stirrer speeds during CSGSR experiments: from 8.000 rpm at feed gas flow 5 liters/ hr N T P until 10.000 rpm at feed 20 liters/hr N T P .

Table II. Temp., °C

Feed Gas, liters/hr NTP

W/F grams cat sec

Results of a mmoles/liter

liters Fixed Bed

CSGSR

480 480 480 480

20.0 10.1 7.9 4.6

560 1102 1410 2424

.297 .297 .297 .297

480 480 480 480

20.0 10.1 7.9 4.6

565 1114 1413 2418

.297 .297 .297 .297

» Vapor pressure of ethylbenzene: 14.04 mm Hg; of styrene: 0 mm Hg.


6.

H E Y N E N AND VAN DER BAAN


CH/CO

= AW/F + B{W/FY

with a standard deviation below 0.01. We did not go into this matter further since the empirical equation was sufficiently accurate for our purpose—i.e., the study of the kinetics of the dehydrogenation reaction. For the fixed bed the concentration c at the outlet was calculated numeri cally by the Runge-Kutta method as a function of W/F, c , k, K , and K and the appropriate values for A and Β by Equation 2b: s

E

0

des

_ k{c

c s c o (AW/F

--_cs -

0

1 + X (c

dW/F

E

0

+ B(W/F)*)/k

eq

j

s

(

2

b

- c ) + K ses s

Values of k, K , and K for different reaction temperatures were deter mined by a numerical search procedure, minimizing the function: s


E

/(*, K , K ) - Σ j E

B

C8m>;

c

" [

(4)

C8>7 2

For the CSGSR a mass balance yields: rW = Fc

8

W F

=

c [ l + KE(CO - cs) + K c ] k(c — c — c cn/k ) 8

8

s

Q

8

csil + Κ φ

β

k[

Co

-

cs -

(5)

s

eq

0

-

Cs) +

csco (AW/F +

#scsl B(W/F) /k ] 2

eq

From Equation 5 the concentration c could be solved as a function of W/F, c , k, Κ , and K , and the values of k, K , and K could be obtained again as described above. This approach is based on the assumption that both reactors are ideal. It is shown below that at least for the CSGSR this assumption is not completely valid. Therefore we have calculated the rate and adsorption con stants with the data from the fixed bed experiments. The experiments in both reactors were done at five different initial con centrations of ethylbenzene, each at three temperature levels and four flow rates as shown in Table I. Results of a characteristic experiment are detailed in Table II. The kinetic constants were computed from the data thus obtained. The experimental conversion data and those calculated with the computed s

g

Έ

0

E

Characteristic Experiment

s

0

Productivities, % Benzene

Toluene

Styrene

#2

Selectivity for Styrene,

0 0.1 0.2 0.6

0.5 0.6 0.8 0.9

18.3 29.1 35.7 51.9

6.8 11.8 13.4 17.6

97.4 98.0 97.4 97.2

0 0.3 0.3 0.6

0.6 0.7 0.8 1.1

15.5 26.7 30.9 42.5

6.0 10.3 11.5 16.0

96.6 96.5 96.5 96.6


76


Table III.

Calculated Constants from Fixed Bed Experiments

Temperature, k, X 10~* K , °C liters/(grams cat sec) liters /mmole E

465 480 495

Π

3.88 5.29 6.99

1.8 1.7 1.9

K , liters/mmole 8

2.7 2.4 2.1

h mmoles/liter q

a

.162 .237 .343

fceqfrom data given by Boundy (17).

β

m mol /1


0 0.156

60

0

0291

0

0.178

0.127

0.628

0

AO

20

1000

2000

Figure 9. Fixed bed experiments at 495°C. Styrene productivity vs. W/F. Continuous lines calculated with constants from Table III.

w/F

gcat.sec/Γ

constants were in good agreement. It appeared, however, that the styrene adsorption constant K could vary by a factor of 2 to 5, without influencing the sum of squares of Equation 4 very much. To narrow the value for K we did experiments with a feed containing 8.55 mm ethylbenzene and 6.1 mm styrene at the same temperature levels and flow rates as chosen for the other experiments in the plug flow reactor. This resulted in the values for k, K , and K given in Table III. With these values both the fixed bed experiments with and without styrene in the feed gas and the CSGSR data could be described well, as is illustrated in Figures 9 and 10 for the experiments at 4 9 5 ° C . Figure 10 shows that the measured styrene productivities for the CSGSR are a little higher than those predicted from the fixed bed experiments. One reason for these deviations is that the selectivity for styrene in the CSGSR is about 1% lower than for the plug flow reactor (see Table II). Another reason is that at least the CSGSR is not ideal. This tends to bring the conversion curves for the plug flow reactor and the CSGSR closer. For the CSGSR the deviation from ideality can be calculated, visualizing the reactor as a recirculated plug flow reactor and obtaining the number of recircu lations from the pressure drop over the bed and the feed flow rate. The data of this calculation are shown in Table IV, and the corrected conversion levels are represented in Figure 10 by dotted lines. From Figure 10 it follows that the s

s

E

s


6.

H E Y N E N A N D V A N DER B A A N


Table IV. Comparison of Styrene Productivities in an Ideal Plug Flow Reactor, an Ideal Mixer, and a Recirculated Plug Flow Reactor at 4 9 5 C e

c mmoles/liter


0

β

Styrene productivities, %

W/F w /r grams cat sec/liter

Ideal Mixer

Ideal Plug Flow

Plug Flow with η Recirculations

.062 .062 .062 .062

5 9 11 17

600 1200 1800 2400

27.1 42.3 52.0 58.8

31.1 52.2 66.5 76.1

27.7 43.2 53.2 59.8

.156 .156 .156 .156

5 9 11 17

600 1200 1800 2400

24.1 38.2 47.4 53.9

27.4 46.7 60.1 69.5

24.6 39.0 48.4 54.8

.291 .291 .291 .291

5 9 11 17

600 1200 1800 2400

20.8 33.5 42.0 48.2

23.3 40.4 52.8 61.8

21.2 34.2 42.9 49.0

.628 .628 .628 .628

5 9 11 17

600 1200 1800 2400

15.5 25.8 32.8 38.2

17.0 30.5 40.5 48.3

15.8 26.3 33.4 38.8

β.99 X 10- so s

H

c

c

c c D F k ke K K , K ; L r u W g m

t

2

a

(l

t

E

concentration of ethylbenzene under reaction conditions, mmole/liter concentrations of styrene and hydrogen, mmoles/liter concentrations of ethylbenzene and styrene under average reaction conditions before entering the catalyst bed, mmoles/liter measured styrene concentration, mmoles/liter concentration of component i in the reactor, mmoles/liter axial dispersion coefficient, m /sec total gas feed to the reactor under reaction conditions, liters/sec reaction rate constant, liters/ ( gram cat sec) equilibrium constant of the dehydrogenation reaction, mmoles/liter adsorption constant of component i, liters/mmole adsorption constant of ethylbenzene and styrene, liters/mmole index number of experimental data at one reaction temperature lengths of the catalyst bed, m reaction rate, mmoles/ (gram cat sec) linear gas velocity in the catalyst bed under reaction conditions, m/sec catalyst weight, grams

s

Literature Cited 1. De Jong, J. G., Batist, Ph. Α., Rec. Trav. Chim. Pays-Bos (1971) 90, 749. 2. Steenhof de Jong, J. G., Guffens, C. H. E., van der Baan, H. S.,J.Catalysis (1972 ) 26, 401. 3. Heynen, H. W. G., van der Baan, H. S.,J.Catalysis, in press. 4. Miller, S. Α., Donaldson, J. W., Chem. Process Eng. (Dec. 1972) 48, 37. 5. Ohlinger, H., Stadelmann, S., Chem.-Ing.-Tech. (1965) 37, 361. 6. Balandin, Α. Α., Tolstopyatova, Α. Α., Zh. Obsch. Khim. (1947) 17, 2182. 7. Carra, S., Forni, L., Ind. Eng. Chem., Process Design Develop. (1965) 4, 28 8. Wenner, R. R., Dybal, E.C.,Chem. Eng. Progr. (1948) 44, 275. 9. Choudhary, V. R., Doraiswamy, L. K., Ind. Eng. Chem., Process Design Develop (1972) 11, 420. 10. Kiperman, S. L., Kinet. Ratal. (1972) 13, 562. 11. Brisk, M. L., Day, R. L., Jones, M., Warren, J. B., Trans. Instn. Chem. Engrs. (1968 ) 46, T3. 12. German, A. L., Heynen, H. W. G.,J.Sci. Instrum. (1972) 5, 413. 13. Hougen, Ο. Α., Ind. Eng. Chem. (1961) 53, 509. 14. Satterfield, C. N., Sherwood, T. H., "The Role of Diffusion in Catalysis," AddisonWesley, Reading, Mass., 1963.


6. HEYNEN AND VAN DER BAAN


15. Balaridin, Α. Α.,Αdvan.Catalysis (1958) 10, 96. 16. Carra, S., Chim. Ind. (Milan) (1963) 8, 949. 17. Boundy, R. H., Boyer, R. F., "Styrene: Its Polymers, Copolymers, and Deriva tives," Reinhold, New York, 1952. 18. Heynen, H. W. G., Thesis, Eindhoven, 1974.


RECEIVED January 2, 1974.


Ethylbenzene Dehydrogenation Kinetics in a Fixed Bed and a

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