An integrated dehydrogenation-hydrogenation membrane reactor for the simultaneous production of styrene and cyclohexane

Reaction Kinetics and the Developmentof Catalytic Processes G.F. Froment and K.C. Waugh (Editors) 1999 Elsevier Science B.V.

229

An i n t e g r a t e d d e h y d r o g e n a t i o n - h y d r o g e n a t i o n m e m b r a n e reactor for the s i m u l t a n e o u s production of styrene and c y c l o h e x a n e T.M. Moustafa-, I. Ashour ~ and S.S. Elnashaie c a Chemical Engineering Department, Faculty of Engineering, Cairo University, Cairo, Egypt b Chemical Engineering Department, Faculty of Engineering, Elmenya University, Egypt c Chemical Engineering and Pilot Plant, National Research Center, Dokki, Cairo, Egypt Abstract A rigorous heterogeneous model is used to study the performance of the membrane catalytic reactor for the dehydrogenation of ethylbenzene to styrene. The mathematical model is extended to simulate a novel hybrid membrane reactor. This reactor is composed of two catalytic sections separated by selective composite membrane for hydrogen separation. One side of the reactor is a dehydrogenation section in which ethylbenzene is dehydrogenated to styrene, while the other catalytic side is a hydrogenation section in which benzene is catalytically converted to cyclohexane. Hydrogen gas is transferred from the dehydrogenation side to the hydrogenation side through the selective composite membrane. The continuous removal of hydrogen from the dehydrogenation section leads to the shift of equilibrium conversion in this section, thus higher styrene yield is obtained. Detailed paprametric investigation has been carried out for the membrane reactor and the hybrid reactor configurations. The effect of co-current and counter-current flow pattern is investigated. Superior performance in terms of ethylbenzene conversion enhancement far above the equilibrium value was observed in the hybrid reactor configuration. The selectivity and styrene yield highly exceed the industrial figures. 1. I N T R O D U C T I O N The use of palladium membranes in laboratory scale catalytic reactors to break the thermodynamic barrier of reversible reactions has been tried successfully by many investigators for a number of important catalytic dehydrogenation reactions [1-4]. However the industrial implementation of this technology requires cheaper and more durable membranes. Recent advances in the manufacture of composite membranes represent an important step forward in the direction of the industrial exploitation of the membrane reactor technology [5-7]. The industrial production of styrene (ST) from ethylbenzene (EB), over iron oxide catalysts and with steam as diluent is a typical industrial example of such reactions where a membrane reactor configuration can improve conversion considerably well above the equilibrium conversion. Most research work suggested the hydrogen permeated to the other side of the membrane to be swept by any inert gas or vacuum [8]. Reaction couphng between the catalyst bed and the sweep gas sides was proposed by a number of investigators [9,10]. If hydrogen permeated from the catalytic dehydrogenation side is utilized in a hydrogenation reaction on

230 the other side of the membrane, a new more efficient configuration can thus be introduced which we will call the Hybrid Reactor (HR). The proposed hybrid configuration in the present work allows the utilization of the permeated hydrogen from the dehydrogenation side to produce cyclohexane through the catalytic hydrogenation of benzene (BZ) on the other membrane side. Any other hydrogenation reaction can be used depending on the industrial situation.Figure (1) gives a schematic diagram of the proposed hybrid reactor. 2. D E V E L O P M E N T O F T H E M O D E L E Q U A T I O N S The following assumptions are used in the derivation of the model: 1. The reaction mixture is an ideal gas in both catalytic reactor sections. 2. Both sections of the reactor are operated at steady state conditions. 3. Radial and axial diffusions are neglected in the two catalytic beds. 4. Concentration profiles are symmetrical around the catalyst pellet center and pellet is isothermal due to high catalyst thermal conductivity. 5. For the catalyst pellets in the dehydrogenation section (DS): The diffusion model used is based on the rigorous Stefan-Maxwell equations (dusty gas model) [ 11-12] with negligible external resistances. 6. For the catalyst pellets in the hydrogenation section: The diffusion inside the pellet is described by Fickian diffusion with negligible mass transfer resistance. 2:1 R a t e E x p r e s s i o n s

i. Catalytic d e h y d r o g e n a t i o n o f e t h y l b e n z e n e Most literature suggested six independent reactions [ 13]: C6H6CH2CH 3 <=>C6H5CHCH 2 + H 2 (:(;H5CH2CH 3 ~ C6H 6 +C2H 4, C6H5CH2CH 3 ~ C6H5CH 3 + CH 4 2 H20 + C2H 4 ~ 2 CO + 4 H 2 , H20 + CH 4 ~ CO + 3 H 2, H20 + CO ~ CO 2 + H 2 The rate equations are given as follows [13]: R1 = kl(PEB - Ps'r P u z / KEB) , R e = ke pEB, R.~ = k:~ pEB pile, R4 = k4 pHeo (pETH) 0"5, R5 = k5 pH20 pMET a n d R6 = k6 ( P r / T'9 pH20 p c o (1-6) The equilibrium constant of the main reaction is given by: KEB -- exp(-AFo/RT) an(! AF o = a + b T + c T 2 (7,8) where a = 122725.16 ; b = -126.27 and c = -2.19x10 s The rates of reaction have the units of kmol/h.kg Cat., and the kinetic constants of the reactions are expressed by: k i = e x p ( A i - E i / R T ) (9) m a i n reaction: s i d e reactions:

The intrinsic rate constants of these reactions are given in Table 1 by [13]. The catalyst for this EB dehydrogenation reaction is iron oxide (Fe20 3) promoted with potassium carbonate (K2CO 3) and chromium oxide (Cr203). ii. C a t a l y t i c h y d r o g e n a t i o n of b e n z e n e On the other side, the catalytic hydrogenation reaction taking place is: C6H 6 + 3H 2 ~ C6H 12 The catalyst is supported nickel and the kinetic

Table 1 F r e q u a n c y factors and Activation energy Reaction Ai Ei 1 0.851 90891 2 14.00 207989 3 0.5(; 91515 4 ().12 103996 5 -3.21 65723 6 21.24 73628

231

rate equation is given by[14]:

kKBPBZ PH2 -- RRBZ = (1 + KBPBz ) (PH, + PBZ )

Where the constants in the rate equation are expressed by the following equations [ 14]: k=k o exp(-E/RT) and KB=KBoexp(EB/RT ) (11,12) Constants for the above equations are given in Table 2 [14].

(10)

Table 2 Constants of benzene kinetic equation

Constant ko E KBo ER

Value 4.36x105 12000 7.88x10 ~ 6000

2.2 M o d e l e q u a t i o n s

For the six reactions in the DS the six material balance equations are given by: dZi/dL=qip~AnRi/FFEB and dXj/dL=qjptr4eRj/FFH2o (13,14) where i represents the reactions 1,2 and 3 and Xi is the fractional conversion of EB in each of these reactions; while j is for reactions 4,5 and 6 and Xj is the fractional conversion of the steam in each of these reactions. The differential mass balance of the hydrogenation reaction is: dXBz2/ dL=rip B:A Be(R') / FFBz 2 (15) The energy balance equations for the two sections are as follows: 1o

dT

6

~" F~Cp - - = ~ (-AH , )rljRjpBA , +U(rcND)(T- T') i=! dL J=i

(16)

i

3 dT' ~'F~Cp ~ + ,-1 ' dL

dOp , , - Cp,, (T - T ) = ( - A H ' ) R qp~ As, - U ( z N D ) ( T - T') dL

(17)

The last term in equations (16) and (17) represent the heat exchange between the reactor sections. The Ergun equation is used for the pressure drop in both reactor sections: r

dP dL

-

10- 5 (1- ~)Go /150pc~(1- ~) [ dp t;3p(;g~ dp + 1.75Go.

(18)

L

The permeation rate of hydrogen through the composite palladium-ceramic membrane is given by [5,6] "

------~--:P dL

". - p ~ . )

(19)

2.3 C a t a l y s t p e l l e t m o d e l The dusty gas model equations for the catalyst particles in the DS are given by [11]: dC i N, ~ Yi N I - Y j N, =--+ z_., " (20) dz D x, ~ g*, D:I

with the boundary conditions: z=O at N~=O a n d z=z o at Ci(zo)-Cis where z is the characteristic length of the pellet. The mass balance equation for the catalyst particles in the HS are given by: d2C, 2 dC, p.,.R' ~ + . . . . (21) dr 2 I" dr D ea

232

de,

with the boundary conditions: Ci=Ci~ at r=rp and--~r =0

at

r=O

In the presence of external heat transfer resistance for the catalyst pellet, but negligble intraparticle heat transfer resistances (which is the case of nickel catalyst), the energy balance equation can be reduced to: Ts = T + ( - A H-' -! ps ! r 2R' dr '

9

(22)

-

h r;

.

The performance of the catalyst pellet in both reactor sections is expressed in terms of the effectiveness factors of each reaction. The effectiveness factors are l

computed using the following well-known relation:

r,

qi = , [ R i dr rpR,, "o

(23)

3. S O L U T I O N O F THE M O D E L E Q U A T I O N S The catalyst pellet equations in both beds are solved using the orthogonal collocation method [15]. The bulk phase differential equationd are integrated using subroutine DGEAR (IMSL Math/PC-Library) [16]. 4. R E S U L T S AND D I S C U S S I O N The data used to simulate the DS is that of the industrial reactor at Polymer Corporation, Sarnio, Ontario, Canada and is given in Table 3. Before exploring the hybrid reactor performance, we will first study the behavior of a catalytic dehydrogenation reactor producing styrene and having membrane with sweep gas in the permeation side.

Table 3 Specifications Item

Bed length Cross sectional area Catalyst density Catalyst diameter Inlet pressure Inlet t~ml)erature Feed: Ethylbenzene Styrene Benzene Toluene Steam Total molar feed

a n d f e e d c o n d i t i o n for D S Value Dimension

1.70 2.98 2146.3 4.7 2.4 922.59 36.87 0.67 0.11 0.88 453.1 491.87

m m2 kcat/m 3 mm Bar K kmol/h kmol/h kmol/h kmol/h kmol/h kmol/h

4.1 M e m b r a n e r e a c t o r f o r t h e p r o d u c t i o n o f s t y r e n e w i t h i n e r t s w e e p g a s The heterogeneous model of the DS is used to simulate this case in which the industrial reactor is fitted with hydrogen selective membrane and the sweep gas is inert and flowing counter currently. Figure (2) shows a considerable EB conversion and ST yield increase over the case without membrane, espicially at higher sweep gas flow rates. For high enough flow rates of sweep gas, EB conversion approaches 52.7% compared with 45.75% for the reactor without membrane. This represents an improvement of 15.2%. The corresponding values of ST yield are 48.2% and 40.17% indicating an improvement of 20%. The same figure depicts the counter-current case. At low flow rates of sweep gas, and using this configuration, ST yield is less than the case without membrane. A typical value is 41.8% at a sweep gas flow rate of 500 kmol/h. At higher gas

233 rates, styrene yield in case of counter-current mode exceeds t h a t of co-current. Values of 48.8% and 48.2% at 3000 kmol/h are observed for the two configurations respectively. The same pattern is observed for EB conversion with values of 53.1% and 52.7% respectively for the same gas rate. Figure (3) shows the hydrogen partial pressure profiles in the reaction and permeation side for a case of counter current mode having a sweep gas flow rate of 100 kmol/h (which is relatively a low flow rate). As shown from the figure, there is a zone of hydrogen back diffusion (i.e. from the permeation side to the reaction side). This explains the reason for the low conversion values. To overcome this region of hydrogen back diffusion, partial blinding of sweep gas can be used. Figure (4) demonstrates the profiles of hydrogen partial pressures in both reaction and permeation sides in case of sweep gas flow rate of 100 kmol/h. The optimum location of partial blinding is at a length of 0.765 m from reactor inlet. EB conversion and ST yield are increased by 2.0% and 6.5% compared with the counter current without partial blinding mode.

4.2 Hybrid R e a c t o r s i m u l a t i o n Data used for simulation of the HR is that present in Tables 3 and 4. Table 5 gives the results of reactor simulation. HR variables are demonstrated versus the reactor length through figures 5 to 8. As shown in figure 5, the trends of EB conversion and ST yield are similar to that of the conventional reactor, but with higher values and higher selectivity towards ST. The same figure shows the profile of the DS temperature. The profile declines with reactor length as the main reaction is endothermic.

Table 4 Specification for HS Item Value and unit Benzene Feed 20 kmol/h Cross sectional area 3.0 m 2 Catalyst diameter 1.86x10 "3 m Catalyst porosity 0.35 Catalyst density 1200 kcat/m 3 Inlet temperature 400 K Inlet pressure 1.1 bar Bed length 1.7 m

Table 5 Results of the HR On the other side of the membrane, and by Item Value and unit the continious transfer of hydrogen to the HS, EB conversion 0.518 benzene conversion increases. The rate of ST yield 0.466 benzene hydrogenation is increased sharply ST production rate 17.855kmol/h after 0.5 m from reactor inlet due to the high D.S exit temperature 841.4 increase in reaction temperature. The other D.S exit pressure 2.36 BZ conversion (H.S) 0.40 part of the figure depicts the temperature of H.S exit temperaure 957.4 the HS of both the bulk gases and the catalyst H.S. exit pressure 1.097 pellets. Small differences between the two temperatures are due to high external heat transfer coefficient (h) between the bulk gas and the catalyst pellet. Figure (8) demonstrates the hydrogen partial pressures in the two reactor sections. Both profiles exhibit a maximum and that of the HS approaches zero after 1.2 m from reactor inlet. This behavior offers a maximum driving force for hydrogen permeation between the two reactor sections.

234

4.3 Effect of use of i n d u s t r i a l m e m b r a n e on reactor p e r f o r m a n c e In order to predict the performance of the reactor on a more practical basis, a commercial membrane was used in the Simulation. Figure (9) shows the profit gain (over the conventional reactor) when using a commercial composite membrane of Bend Inc. in the proposed hybrid reactor. Membrane area was varried to predict the range of ecconomical optimum design. Calculations were based on membrane cost of $300 /ft 2 [17]. The increase in styrene yield was calaulated with a MT price of $660. Rate of return on investment was calculated based on the profit gained with respect to the additional investment (membrane cost). The optimum membrane area can be taken at the point of maximum rate of return. 4.4 Possible process design i m p r o v e m e n t s o f styrene p r o d u c t i o n p l a n t s The above simulation results indicate some of the benefits associated with the use of the HR configuration. The benefits are not limited to the increase in ST yield and selectivity but also includes the useful utilization of the permeated hydrogen to produce cyclohexane. The improvements in the design of such reactors can allow for the complete elimination of the recycle step for the unreacted EB. Exchange of heat between the two beds is assumed to take place espicially that new approaches of membrane preparation on highly conductive substrates are now feasible. Figure 9.a demonstrates the profiles of EB conversion and ST yield, reaching values of 93 and 87% respectively. This value of ST yield represents an enhancement of 116% over the industrial conventional reactor. Figure 9.b depicts the temperature profile in the DS, which shows a minimum. This behavior is due to the competing effect between heat lost in the endothermic dehydrogenation reaction and heat gain by the transfer of heat from the HS. 5. CONCLUSIONS Hybrid configuration improves the performance considerably, namely increasing EB conversion, ST yield and selectivity. The magnitudes of the increases are quite appreciable together with the production of additional product, which is cyclohexane. A design for the HR was proposed with longer length compared with the industrial case, as thermodynamic barrier has been broken, and was found to give values of styrene yield which is 116% more than the industrial reactor. The quantitative predictions obtained using this rigorous model should be reasonably reliable, however, they still need to be checked experimentally. 6. R E F E R E N C E S 1 2 3 4 5 6 7 8 9

N. Itoh, AICHE, 33 (1987) 1576. N. Itoh, J. Chem. Eng. Jpn., 24 (1991) 664. Y.V. Gokhale, R.D. Noble and J.L. Falconer, J. Memb. Sci., 77 (1993) 197. E. Gobina, K. Hou and R. Hughes, J. Memb. Sci., 105 (1995) 163. R. Govind and D. Atnoor, Ind. Eng. Chem. Res., 30 (1991) 591. J.P. Collins and J.D. Way, Ind. Eng. Chem. Res., 32 (1993) 3006. K.L. Yeung and A. Varma, AICHE J., 41 (1995) 2131. B.K. Abdalla and S.S. Elnashaie, J. Membr. Sci., 85 (1993) 229. N. Itoh, J. Chem. Eng. Jpn., 23 (1990) 81.

235 10 E. Gobina and R. Highes, Chem. Engng. Sci., 51 (1996) 3045. 11 E.A.Mason and A.P.Malinauskas, Elsevier, Amsterdam 1983. 12 S.S.Elnashaie, B.Abdalla and R.Hughes, Ind.Eng.Chem.Res.,32(1993) 2537. 13 J.G.P. Sheel and C.M. Crowe, Can. J. Chem. Eng., 47 (1969) 183. 14 J.K. Marangozis, B.G. Mantzouranis and A.N. Sophos, Ind. Eng. Chem. Prod. Res. Dev., 18 (1979) 61. 15 J. Villadsen and M.L. Michelsen, Prentice Hall, New York, NY, 1987. 16 A.C. Hindmarsh, Gear, Lawrence Livermore Lab., 1974. 17 Bend Research Inc., (1996).

236

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