Dehydrogenation of butene-1: effect of coking on reactor performance

applied catalysis A ELSEV I ER

Applied Catalysis A: General 119 (1994) 153-162

Dehydrogenation of butene- 1" effect of coking on reactor performance R. H u g h e s *, C . L . K o o n Department of Chemwal Engineering, Universtty of Salford, Salford M5 4WT, UK Received 29 April 1994; revised 14 June 1994; accepted 14 June 1994

Abstract The dehydrogenation of butene-1 over a chromia-alumina catalyst has been used to study the distribution and concentration of the coke deposited along a fixed bed reactor using different feed concentrations of butene-1. Axial temperature measurements showed a steep initial temperature minimum followed by a subsequent shallower minimum. Corresponding coke profiles were determined using a microbalance and confirmed by the non-invasive neutron attenuation technique. These together with determinations of exit gas concentrations suggest that coke is formed primarily from the reaction product butadiene, and that this coke deposit limits further formation of this product and thus decreases further coke formation. Keywords: Butene dehydrogenation; Chromia-alumina; Coke deposition

1. Introduction Most butadiene production is obtained by catalytic dehydrogenation of n-butane or n-butenes or a mixture of these gases. The reaction is endothermic and produces a carbonaceous deposit which deactivates the catalyst. To minimise both these influences, the feed is normally diluted with steam or operated under reduced pressure. Nevertheless, deactivation by these coke deposits cannot be eliminated entirely, so that the reaction process in a single reactor is necessarily periodic, with the dehydrogenation reaction stopped after a certain time when the coke on the catalyst is then burnt off by oxidation [ 1 ]. Because of this requirement and to obtain a continuous production the process is generally operated using a multiple reactor system. * Corresponding author. E-mail [email protected], tel. ( + 44-61 ) 7455081, fax. ( + 44-61 ) 7455999. 0926-860X/94/$07.00 © 1994 Elsevier Science B.V. All rights reserved SSDI 0926-860 X ( 94 ) O0139-I

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The kinetics of the process are complex with simultaneous formation of the butadiene product and the coke deposit, but measurements of the kinetics have been made by Forni et al. [2], Happel et al. [3], Timoshenko and Bayanov [4] and more recently by Mandani [ 5 ]. However, the most authoritative studies have been conducted by Froment and co-workers [ 6-8 ] who undertook comprehensive investigations of both the dehydrogenation reaction and the rate of coke formation and analyzed their results using Langmuir-Hinshelwood rate expressions. One area of uncertainty in previous work is the nature of the main coke precursor. From their results on the dehydrogenation of n-butene to 1:3 butadiene, Dumez and Froment [ 6 ] concluded that coke was formed from both butenes and butadiene. Subsequently, Acharya and Hughes [9] used the kinetics established by Dumez and Froment in a simulation model of n-butene dehydrogenation for both catalyst pellets and a fixed bed reactor. Under the conditions used in their analysis coke deposition by a series mechanism i.e. from the produced butadiene, was shown to be the preferred mechanism. The present paper reports results on the deactivation arising from coke deposits during the dehydrogenation of butene-1 in integral reactors using a commercial chromia-alumina catalyst in pellet form. Temperature profiles were measured along the axis of the reactor as well as analysis of the exit concentration of the gaseous products. Coke profiles were determined using conventional gravimetric analysis and also by the neutron attenuation technique.

2. Experimental Dehydrogenation of butene-1 was carried out in a fixed bed reactor of length 1.14 m and 25 mm in diameter. The central 0.86 m of this reactor was packed with 4 mm equant pellets of a 19% Cr203/alumina catalyst. A preheater heated the inlet gases and three separate windings along the main reactor enabled a uniform initial overall bed temperature within + 5°C to be attained. During the heating period to the reaction temperature a flow of nitrogen was passed through the reactor and this was maintained whilst at the reaction temperature until a stable uniform temperature profile was obtained. The flow was then changed to the butene-1/nitrogen mixture composition required. Temperature measurements along the axis of the bed were made with six 0.5 mm diameter chromel-alumel thermocouples placed at intervals of 0.14 m. Feed gas mixtures containing 20, 30 and 40% butene-1 in nitrogen were set at the required flow-rates (usually 1.67.10- 5 m3/s) using mass flow controllers. The initial reactor temperature in all experiments was set at 450°C. Exit gas concentrations were measured using a gas chromatograph. The first gas analysis was carded out after 5 min and subsequent samples at intervals of 20 min until the end of the experiment. Coke profiles were determined by two methods; in the first individual catalyst pellets were taken from different axial points in the reactor and combusted in air using a microbalance. The second method was the

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155

non-invasive method in which the attenuation of a neutron beam served to determine the coke loading at different points within the bed [10]. Satisfactory agreement was obtained between the two techniques. A few experiments were also made using a much reduced catalyst charge in the reactor of 5 g. This charge of the same 4 mm equant pellets was located in the mid section of the reactor and the temperature within this very short catalyst section (ca. 1 cm length) was measured using a single chromel-alumel thermocouple of 0.5 mm O.D. Measurements of the exit gas concentrations were made at two flowrates of 1.67.10 -5 m3/s and 8.35.10 -6 m3/s for a 30% butene-1 feed.

3. R e s u l t s

Measurements of axial temperature profiles, coke profiles and exit gas compositions were measured for the 0.86 m reactor at a total feed flow-rate of 1.67.10 -5 m3/s for the three butene-1 concentrations of 20, 30 and 40%.

3.1. Temperature profiles These were determined for times on stream of up to 6 h and representative examples of the profiles obtained for butene-1 feed concentrations of 20% and 40% are shown in Figs. 1 and 2, respectively. The 30% butene-1 feed gave similar profiles, but intermediate in value between those in Figs. 1 and 2. These figures demonstrate the rapid development of a temperature minimum in the inlet region of the reactor within a short period of time and which was located at a distance of about 0.09 to 0.1 m from the bed entrance. A further more shallow minimum appeared further along the bed at about 0.5 m from the front of the bed. With 0

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R. Hughes, C.L. Koon /Applied Catalysis A: General 119 (1994) 153-162

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increasing time, the first minimum showed a maximum decrease in temperature of 56°C (20% butene-1 feed) to 91°C (40% butene-1 feed). The temperature minimum near the reactor inlet deepened rapidly during the first hour of reaction, whereas the second temperature minimum near the outlet tended to flatten as the time-on-stream was increased. Clearly, increase in n-butene-1 concentration gives a much larger decrease in overall temperature in the reactor.

3.2. Coke profiles Coke profiles for the two feed concentrations of 20% and 40% are given in Figs. 3 and 4, respectively. In both cases little or no coke was found in the inlet region of the bed. Also two maxima in the coke profiles occur. Overall coke levels increase 2.5

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with time-on-stream and feed concentration as expected but rates of deposition tend to decrease after 4 h on stream (Fig. 4). Good agreement was obtained between results obtained from direct sampling of the catalyst pellets in the bed after reaction (as shown) and the neutron attenuation technique; the latter technique, being noninvasive, demonstrated that sampling of the catalyst pellets for microbalance determinations caused no significant error. A direct comparison of the coke and temperature profiles for the 40% butene-1 feed at three different times is shown in Figs. 5, 6 and 7. It can be seen that deposition of coke effectively commences after the first temperature minimum, then declines at about the mid-point of the reactor. Subsequently a second maximum in the coke profile may occur which coincides with the increase in temperature in this region near the reactor exit.

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3.3. Gas exit compositions Gas analysis was performed on the exit gas at frequent time intervals for all three butene-1 feed concentrations. A representative plot is shown in Fig. 8 for the 30% butene-1 case. It can be seen that under the experimental conditions used, not all the butene-1 is converted and that the main product of the reaction is cis-2-butene, followed by butadiene with trans-2-butene giving the lowest product concentration. All products increased rapidly initially but then remained constant after 200 min time-on-stream while after this same time period the conversion of butene-I was constant at 93.3%. Hydrogen in the exit gas was not determined; the amount of this was estimated from the amount of butadiene formed, since by stoichiometry the quantities of

R. Hughes, C.L. Koon /Applied Catalysis A: General 119 (1994)153-162 12

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butadiene and hydrogen should be equal if no further reactions occur. Mass balances with this assumption gave 96% and 94% accountable levels for the 20% and 30% butene-1 feeds and a value of 92% for the 40% butene-1 feed. The estimated hydrogen levels, however, are probably low because some of the butadiene will have been converted to coke. Since the coke content increases with increasing feed concentration of butene-1, this could explain the trend in the mass balance values with increasing butene-1 feed levels.

3.4. Measurements using the 5 g catalyst bed Product gas analysis and the temperature time record for this reactor for a 30% butene-1 concentration in the feed gas and a total flow of 8 . 3 5 . 1 0 - 6 m3/s are given in Fig. 9. No butadiene was detected in this experiment or in a similar experiment at the same butene-1 concentration but with an increased flow-rate of 1.67.10-5 m3/s. Both 2-butenes were present, but now the trans-isomer was predominant. The butene- 1 concentration in Fig. 9, after a small initial decrease, is seen to increase with time to give a final conversion of 40% after 5 h, whereas the butene-2 isomers both decrease with time after a gradual initial increase. These results suggest that severe deactivation of the catalyst bed has occurred and this was confirmed when the catalyst bed was analyzed and found to contain 2.74% coke after the 5 h experiment duration. The low conversion of butene-1 of 40% contrasts with the behaviour of the 0.86 m reactor where a conversion greater than 90% was attained. It was hoped initially, that the small amount of catalyst used would enable the reactor to operate in the differential mode and thus enable some kinetic data to be extracted. However, the high conversion of 40% precluded this, as did an observed temperature drop of 33°C after 100 min. The latter value was approximately one half of the value of the initial minimum in Fig. 1 which varied from - 64°C at 2 h to - 67.5°C at 4 h.

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R. Hughes, C.L. Koon/Apphed Catalysts A: General 119 (1994) 153-162 20

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4. Discussion The results from the 0.86 m length reactor show some interesting features. The reactor was heated throughout the experiment and was not adiabatic in operation. The initial temperature drop of 56-91°C was attained after 6 h operation, but even after 1 h a temperature decrease of - 40 to - 80°C was observed, the more rapid decrease corresponding with an increase in butene-1 concentration. This tempera° ture minimum must be caused by dehydrogenation of the butene-1 since only butadiene production gives sufficient endothermicity (both isomerisation reactions are slightly exothermic at the temperature of operation). The butadiene formed does not produce much coke, presumably due to the low temperature created by the endothermic process. The decrease in temperature also reduces the overall reaction rate; the temperature will then increase in the bed due to heat from the reactor windings. Under these conditions the butadiene produced in the inlet region will probably be the main cause of the coke deposition which now becomes abundant. The coke production will cause some deactivation of the bed at this point but as the temperature has now increased more butadiene will be produced, but to a lesser extent than previously due to the lower butene-1 concentration and some deactivation of the catalyst bed. This will cause the second shallow minimum to form. Finally the bed tends to return to its original temperature due to decrease in reactant concentration and catalyst deactivation. The exit gas concentrations do not give butadiene as the major product due to the comparatively low initial temperature of 450°C, compared to the more favoured temperature for butadiene production of 580 to 600°C. Thus, comparable amounts of the butene-2 isomers are produced in the present system. This is confirmed by a comparison of selectivities, these for butadiene obtained in the present work being

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in the range of 22 to 25%, whereas in commercial operation at higher temperatures typical selectivities are in the range of 77 to 86%. When only 5 g of catalyst were employed in the reactor, the results obtained were unusual. Although the temperature as measured at the centre of the catalyst bed decreased with time the overall temperature decrease was much smaller than in the longer reactor. The fact that a temperature drop was observed suggests that butadiene should be formed, yet no butadiene was detected in the exit gas. However the coke level in the bed is relatively high at 2.47% after 3 h at a standard flow-rate of 1.67.10-5 m 3 /s. A possible explanation is that the butadiene formed produces coke with production of hydrogen. The latter then reacts with any remaining butadiene to give the butene-2 isomers. The effect of hydrogen produced has been noted previously by Pena et al. [ 11 ]. The space velocities for the two reactors are very different; that for the 0.86 m length reactor being 0.045 s - 1 whilst that for the 5 g catalyst bed is 3.34 s - 1. The 5 g catalyst bed is probably not fully equilibrated and this could account for the change in yield of the two butene-2 isomers with the trans-isomer now being larger in contrast to the 0.86 m reactor where the cis-isomer was produced in greater yield. Equilibrium calculations on the butene-2 isomers predict that the trans-isomer should be predominant within the temperature range of the present experiments. This was observed to be the case for the 0.01 m length reactor and has been confirmed for the initial stages of reaction in a differentially operated reactor by Pena et al. [ 11 ]. However, at longer reaction times the latter authors noted that the concentrations of the two isomers became almost equal. The observed reversal of the levels of the butene-2 isomers obtained with the 0.86 m length reactor can only be attributed to subsequent reactions, possibly coupled with the coke distribution, which occur along the much greater length of this reactor. The continued production of 1:3 butadiene with time in the 0.86 m length reactor is surprising, but it should be noted that the concentration of the butene-1 feed is also decreasing with time. If overall deactivation was occurring then the opposite behaviour would be expected for both components. The fact that this does not occur in this reactor, compared with the behaviour in the 0.01 m length reactor where orthodox concentration/time trends consistent with severe deactivation occur, once again suggests that the results obtained with the longer reactor depict the complexity of reactions in this system.

Acknowledgements We gratefully acknowledge the financial support of the Science and Engineering Research Council (UK) for this research.

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References [ 1 ] C.L. Thomas, Catalytic Processes and Proven Catalysts, Academic Press, New York, 1970. [2] L. Forni, L. Zandrighi, S. Carra and C. Canenaghi, J. Catal., 15 (1969) 153. [3] J. Happel, H. Blank and T.D. Hamil, Ind. Eng. Chem. Fundam., 5 (1966) 3. [4] V. Timoshenko and R.A. Bayanov, Int. Eng. Chem., 2 (1972) 314. [5] F.M. Mandani, Ph.D. Thesis, Salford University, 1991. [6] F.J. Dumez and G.F. Froment, Ind. Eng. Chem. Process Des. Dev., 15 (1976) 291. [7] J.W. Beeckman and G.F. Froment, Ind. Eng. Chem. Fundam., 18 (1979) 245. [8] G.B. Matin, J.W. Beeckman and G.F. Froment, J. Catal., 97 (1986) 416. [9] D.R. Acharya and R. Hughes, Can. J. Chem. Eng., 68 (1990) 89. [10] C.L. Koon, D.R. Acharya and R. Hughes, J. Catal., 126 (1990) 306. [ 11 ] J.A. Pena, A. Monzon and J. Santamaria. J. Catal., 142 (1993) 59.