Mathematical simulation of catalytic dehydrogenation of ethylbenzene to styrene in a composite palladium membrane reactor

Mathematical simulation of catalytic dehydrogenation of ethylbenzene to styrene in a composite palladium membrane reactor

journal of MEMBRANE SCIENCE ELSEVIER Journal of Membrane Science 136 (1997) 161-172 Mathematical simulation of catalytic dehydrogenation of ethylben...

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journal of MEMBRANE SCIENCE ELSEVIER

Journal of Membrane Science 136 (1997) 161-172

Mathematical simulation of catalytic dehydrogenation of ethylbenzene to styrene in a composite palladium membrane reactor Ch. Hermann,

E Quicker, R. Dittmeyer*

Lehrstuhl fiir Technische Chemie 1, Universitiit Erlangen-Niirnberg, Egerlandstrafle 3, D-91058 Erlangen, Germany Received 10 December 1996; received in revised form 19 June 1997; accepted 23 June 1997

Abstract The catalytic dehydrogenation of ethylbenzene to styrene was studied in a tubular palladium membrane reactor using a commercial styrene catalyst. A mathematical model of the membrane reactor is presented which takes into account the different mass transport mechanisms prevailing in the various layers of the membrane, that is, multicomponent diffusion in the stagnant gas films on both faces of the membrane, combined effective multicomponent diffusion, effective Knudsen diffusion, and viscous flow in the macroporous support, Knudsen diffusion in the microporous intermediate layer, and Sieverts' law of hydrogen transport through the Pd-film. A kinetic model from the literature [1-3] was adjusted to match conversion and selectivity observed during experiments with a commercial catalyst in a laboratory fixed-bed reactor. Simulation calculations based on the resulting effective kinetics were carried out for industrially relevant operating conditions and various process configurations, that is, use of inert sweep gas, evacuation of the permeate gas, and oxidation of the permeated hydrogen. The results demonstrate that with the present catalyst used under typical process conditions (T, P, WHSV, S/O) removal of hydrogen through the membrane gives only a small increase of styrene yield. However, the model predicts that by increasing the reaction pressure in the membrane reactor the kinetic limitation can be overcome and ethylbenzene conversion can be increased to above 90% without markedly decreasing styrene selectivity.

Keywords: Ethylbenzene dehydrogenation; Membrane reactor; Composite membranes; Palladium membranes; Mathematical model

1. Introduction Styrene is one of the most important monomers for the manufacture of thermoplastics. Worldwide styrene demand is forecast to exceed 20 MMton by the year 2000 at an average annual growth rate of 5 to 6% [4]. More than 90% of total styrene production is based on catalytic dehydrogenation of ethylbenzene, operated *Corresponding author. Tel.: +49-9131-857431; fax: +49-9131857421; e-mail: [email protected] 0376-7388/97/$17.00 @ 1997 Elsevier Science B.V. All rights reserved. PII S0376-73 88(97)00 160-9

at 550-650°C and atmospheric or sub-atmospheric pressure over Fe203-catalysts doped with K2CO3 and several metal oxides, for example, Cr203. Industrially, the reaction is carried out in the presence of large amounts of steam, the molar steam-to-oil ratio S / O typically ranging from 6 to 12. Ethylbenzene conversion reaches up to 70%; the selectivity to styrene is around 95%. Fig. 1 shows a simplified reaction scheme of catalytic ethylbenzene dehydrogenation. The main reaction to styrene and hydrogen is reversible; benzene and toluene are the dominating

Ch. Hermann et al./Journal of Membrane Science 136 (1997) 161-172

162

(+ H2)

Ethylbenzene

- H2 ~

Benzene + C2H 4

+ H2 ~ Styrene

'

Toluene + C H 4 + H20

Coke

+H20

, CO/CO 2

Pore size dp [urn] 0.005 0.1 1

Fig. 1. Simplified reaction scheme o f catalytic d e h y d r o g e n a t i o n of ethylbenzene.

side products accompanied by ethylene and methane. Additional side products are coke and carbon oxides, the latter being generated by the gasification of coke or coke precursors or by steam reforming of methane and ethylene. To increase ethylbenzene conversion beyond the thermodynamic equilibrium several authors have suggested the use of permselective membranes in order to remove hydrogen from the reaction mixture and hence to suppress the reverse reaction [1,5-15]. For high temperature dehydrogenation reactions three different types of membranes have been proposed: dense metal membranes (e.g. Pd or Pd-alloys), porous membranes (e.g. alumina, titania, zirconia, or vycor glass), and composite membranes (e.g. metal/alumina, metal/ vycor glass, or metal/stainless steel, where metal most frequently is Pd or one of its alloys) [16-18]. Dense palladium membranes offer the advantage of being permeable only for hydrogen, but suffer from high cost and low permeability, both due to the high wall thickness of 100-150~tm. Porous membranes exhibit sufficient permeability and are available at moderate cost; however, they have a poor hydrogen selectivity, as the smallest available pore size of 4 5 nm is yet too large for molecular sieving. Instead, the separation behavior is governed by Knudsen diffusion, leading to a significant loss of reactants. In addition, the sealing of ceramic or vycor glass to conventional reactor installations (metal) is a serious problem. Composite membranes (see Fig. 2) are used to combine the advantages of the two other types, that is, the high permselectivity of palladium (or other metals) and the high permeability of the porous support. From a practical point of view, porous stainless steel appears more promising than ceramics or vycor glass because the sealing is possible with conventional technology and the structural properties are clearly superior [19].

Fig. 2. Schematics o f a composite m e t a l / c e r a m i c m e m b r a n e .

The driving force for the removal of hydrogen from the catalyst bed is a gradient of the chemical potential of hydrogen between retentate and permeate. Different options exist to establish this gradient, i.e. use of inert sweep gas, application of a pressure differential, or oxidation of the permeated hydrogen, for example, with air (reactive sweep gas) [20]. All these methods decrease the partial pressure of hydrogen on the permeate side to establish a hydrogen partial pressure gradient across the membrane. From an industrial point of view, the use of inert sweep gas is not a viable option because of the large amount of expensive sweep gas required. Besides, a diluted hydrogen stream is produced which can be used only as lowvalue fuel. Applying a pressure differential is well suited for processes operated at elevated pressure. Unfortunately, ethylbenzene dehydrogenation is carried out at low pressure, so that further reduction of the permeate pressure would entail high vacuum costs. On the other hand, with non-porous membranes marketable hydrogen could be produced which could (partly) compensate for these costs. Oxidation of the permeated hydrogen with air is also viable only for nonporous membranes, as otherwise oxygen penetrates into the catalyst bed and decreases the selectivity by forming deep oxidation products. On a whole oxidation is the most attractive option, because neither expensive inert sweep gas nor vacuum are required. Moreover, the exothermic oxidation can be coupled to the endothermic dehydrogenation which leads to better heat utilization and to reduced steam demand.

163

Ch. Hermann et al./Journal of Membrane Science 136 (1997) 161-172

In most of the previous studies on the application of membrane reactors to catalytic dehydrogenation of ethylbenzene commercial ceramic ultrafiltration membranes have been employed [7-14]. Abdalla and Elnashaie [1] in 1994 were the first to report simulation results for ethylbenzene dehydrogenation in a composite Pd/ceramic membrane reactor, followed by Gobina et al. in 1995 [15]. The aim of the present work is to explore the potential benefits of composite Pd-based membranes for ethylbenzene dehydrogenation in more detail and to determine the crucial parameters controlling the overall performance of the membrane reactor.

2. M a t h e m a t i c a l

Retentate

Membrane JH2

"

,,

Permeate ,, I

i

1

"

l

r Fig. 4. Model membrane geometry.

program supports different process options, that is, the use of an inert sweep gas, the evacuation of the permeate gas, or the oxidation of the permeated hydrogen, as well as isothermal or adiabatic operation.

model

2.1. Transport in axial direction A computer program was developed for steadystate tubular membrane reactors consisting of a conventional fixed-bed of catalyst surrounded by an inert membrane. The location of the catalyst may be either on the tube side or on the shell side (see Fig. 3). A pseudo-homogeneous model assuming plug flow of the bulk gas phase for the tube and shell side is used to describe the catalyst bed; radial temperature gradients across the membrane are neglected. Composite membranes are treated by accounting for up to three membrane layers, that is, the macroporous support, a microporous skin layer, and an impervious metal film coated either on the tube side or on the shell side surface. In addition, stagnant gas films on both faces of the membrane are considered (see Fig. 4). The heat transfer resistance through the membrane is neglected, i.e. the permeate gas phase and the retentate gas phase are assumed to be at the same temperature. The

b) tube-side

a) shell-side

configuration

configuration

At the retentate (reaction) side the downstream change of the moles of species i is controlled by the reaction kinetics and by the transport rate through the membrane, as given by: dh/R

dz

m

= ARPb Z

vijrj

-

-

27rRRJi[R,

At the permeate (sweep) side the same type of equation holds, but in the inert sweep gas mode and in the permeate evacuation mode the reaction term is zero. This leads to:

dhSi = 27rRRJiIR,

(2a)

dz For the case of permeate oxidation it is assumed that a given fraction f of the permeated hydrogen is converted to water without explicitly considering the kinetics of hydrogen oxidation. Hence, we have: dh s

dz - 27rRRJiiR. + fi27VRRJH2JR. Fixed-bed

(1)

j=l

(2b)

Fixed-bed

where:

Permeate ~

Retentate ~

Rete~te Fig. 3. Schematics of the reactor configuration.

Perlmea~

j5 = - f j~=f fi=-0.5f j5 = 0

for for for for

i -= H2 i=H20 i=O2 i ~ H2, H20, 02

Under isothermal conditions the temperature gradient dT/dz is zero, whereas under adiabatic

Ch. Hermann et al./Journal of Membrane Science 136 (1997) 161-172

164

conditions dT/dz can be calculated from:

~- =

-(ZXHR~),) + (-ZXHox)fJH2

=~

ntotCp

(3)

The pressure gradient dP/dz along the catalyst bed is calculated according to the Ergun equation [10]:

dpR

- 1 5 0 qu (1

-

(1 -- eb)

£b) 2

whereas for determining the pressure gradient on the sweep gas side the following distinction is made: (a) tube side sweep (Hagen-Poiseuille law)

dP s dz

- -

dz

-

R2

Mass transport in the radial direction involves different transport mechanisms for the various regions of the membrane (cf. Fig. 2). For describing multicomponent diffusion through the stagnant gas films on both faces of the membrane the Stefan-Maxwell equations dyi _ 9¢T k ~-, JjYi - YjJi

pk ~_ 1

k = R, S

(6)

D~

are used [22], where indices R and S denote reaction side and sweep side, respectively. Temperature and pressure are supposed to be constant within the gas film; the film thickness is estimated from Nu-number correlations [23]. To treat effective multicomponent diffusion, effective Knudsen diffusion, and viscous flow through the macroporous support the dusty gas model

dr

(9)

Knudsen diffusion is assumed to control the flux of the various species through the microporous layer [9]. Hence, we have for the flux of species i:

gCTR ( ~ J j Y i - Y j J i P

(l O)

Finally, Sieverts's law is used to describe the hydrogen transport through the impervious metal film [1,25]. Consequently, the hydrogen flux reads: r

2.2. Transport in the radial direction

dyi

DiK = ; 3 - V ~

8z/uln(RR -- d~/Rs)/{ln(RR + d~/Rs)

• [RZs+ (RR + da) 2] + [R2 - (Re + da)2]} (5b)

dr

(8)

(5a)

(b) shell side sweep (modified Hagen-Poiseuille law [10]) de s

B0 - e a 2 r32

e dci Ji ~- --OiK d~

-8qu -

is applied, with parameters Bo and DeK related to porosity, tortuosity, and pore size of the support. According to [24] it holds:

j=l

(e/r)D~0

~e_~ DiK/

with parameter D related to the diffusivity of hydrogen, DH, and the concentration of dissolved hydrogen, Co, in the metal. After Ref. [1] it holds: D --

DHCo ln(ro/ri)v~

(12)

where DH and Co for pure palladium are determined as a function of temperature according to: DH = 2.30 × 10 -7 • e (-21700/9~T)

(13)

Co = 3.03 × 102 • T -1°35s

(14)

A gear-algorithm (NAG-Subroutine D02EBF [26]) is used to integrate Eqs. (1)-(5) starting from the initial conditions: h/klz=0=h~'0; Tklz=0=T k'°

i= 1...n; k=S,R (15)

For this, the molar fluxes through the membrane, Ji, must be specified• We assume that the membrane is inert and operated in steady state; hence, Ji is constant through the various layers. Taking this condition together with the boundary conditions for the concentration of the various species in the retentate and permeate bulk gas phase (cf. Fig. 4):

R.

Cilr=O = Cig, Cilr=cSR+d,+d2+d3+ds = c - Yi ~ +

dr

(7)

s

i = 1... n (16)

165

Ch. Hermann et al./Journal of Membrane Science 136 (1997) 161-172

after discretization, allows to solve the ODE system defined by Eqs. (6)-(14) for Ji with the help of a generalized Newton-Raphson method [27].

I

~

z

~

2

X

×

~

2.3. K i n e t i c m o d e l

The kinetic model was taken from the literature [1-3,21], but the individual rate constants were multiplied by empirical tuning factors to match conversion and selectivity observed in fixed-bed reactor experiments with a commercial 3 mm styrene catalyst provided by Stid-Chemie AG, Miinchen/Germany. Note that the estimates of the kinetic constants determined from these experiments represent the effective k i n e t i c s rather than the intrinsic kinetics, as the influence of diffusion has not been considered in terms of a rigorous heterogeneous modeling, but lumped into the rate constants according to a pseudo-homogeneous approach. Table 1 lists the resulting effective kinetic parameters together with the equilibrium constant of the main reaction and the enthalpy data used for the simulation [29].

I

I

I

7

X

7

X

x

i

i

T

"d

-6

"~

~

ca

B

x

t'¢3

i

% x

x

x

~ o~ I

o

o

0

3. S i m u l a t i o n

©

results ~

The model was used to simulate a membrane reactor equipped with a Pd-coated porous stainless steel tube of 25 mm inner diameter, 2 m m wall thickness, and 1 m length. The catalyst is assumed to be loaded on the tube side (see Fig. 3(b)). The membrane tube is surrounded by a 38 mm i.d. steel tube, creating a shell side annulus of about 5 m m width. Both tube and shell side have roughly the same free cross-sectional area. The pressure drop over the catalyst bed, as calculated according to Eq. (4), was in the range of 0.05 to 0.2 bar, whereas the pressure drop on the permeate side was found to be negligible (Eq. (5b)). Table 2 provides an overview of membrane properties, catalyst properties, and operating conditions.

-

%

% ~-

~v

x

x

×

×

X

X

.~ ~v

II

II

li

il

II

ii

[.-.~

u

~

N

p No ~ '

£ [.-.N

3.1• I n e r t s w e e p g a s ~"

o

Fig. 5 shows the calculated conversion of ethylbenzene as a function of the normalized sweep gas flow and the retentate pressure (i.e. the reaction pressure); the permeate pressure is kept constant at 1 bar. Increasing the sweep gas flow at 1 bar reaction

+

~ #

0 ~ +T

+ 0 rj

+ ~:~

o

~

+

m m

+

+

+

+

m

0

0

0

II

x

II

166

Ch. Herrnann et al./Journal of Membrane Science 136 (1997) 161-172

Table 2 Case description:Composite Pd/stainless steel membrane $tyNme

Membrane properties Porosity of the support Pore size of the support Tortuosity of the support Thickness of the support Thickness of the Pd-layer

0.5 0.2 gm 3 2 mm 10 gm

Catalyst properties Pellet size Fixed-bed void fraction Pellet density

3 mm 0.5 2150 kg/m 3

Operating conditions Temperature WHSV S/O (molar)

620°C 1.0 h- l 12

t

Fig. 6. Pd-coated porous stainless steel. Case I: Use of an inert sweep gas. Styreneselectivityversus sweep gas flow and retentate pressure (T = 620°C, Pe = i bar, WHSV = 1.0 h -1, S/O = 12).

I

I M

I m ~

t

0

[-I

o

Fig. 5. Pd-coated porous stainless steel. Case I: Use of an inert sweep gas. Ethylbenzeneconversionversus sweep gas flow and retentate pressure (T=620°C, Pp= l bar, WHSV=I.0h -l, S/O = 12).

pressure does not markedly increase ethylbenzene conversion. This is, however, not true for higher reaction pressure, for example, at 2.5 bar ethylbenzene conversion reaches about 90% when increasing the sweep gas flow from zero to 100 times the ethylbenzene flow. Note that the sweep gas efficiency decreases with increasing sweep gas flow. Without sweep gas, which marks the fixed-bed mode, an increase of conversion from 60.2% to 74.2% is predicted upon increasing the reaction pressure from 1 to 2.5 bar. The corresponding equilibrium conversion related to the main reaction, as determined by Aspen-Plus calculations, is 84.1% at 1 bar and 71.2% at 2.5 bar. Hence, we note that the reaction at 1 bar does not reach thermodynamic equilibrium,

but instead, is controlled by the kinetics. An increase of pressure leads to a shift of the equilibrium towards lower conversion but also to an increased reaction rate; hence, it drives the reaction closer towards the equilibrium. The fact that the predicted conversion at 2.5 bar reaction pressure exceeds the equilibrium conversion of the main reaction can be ascribed to a significant contribution of the irreversible parallel side reactions to benzene and toluene (reactions 2 and 3 of Table 1). Fig. 6 shows the behavior of styrene selectivity for the same conditions. Again, at 1 bar there is no substantial change of selectivity upon increasing the sweep gas flow. If the reaction pressure is increased in the fixed-bed mode from 1 to 2.5 bar, then styrene selectivity decreases from 95.6% to 83.3% as a matter of the acceleration of the reverse reaction (reaction 1of Table 1) as well as of the parallel side reaction to toluene (reaction 3 of Table 1); both are second-order reactions. However, in the membrane reactor mode increasing the sweep gas flow prevents this drop of selectivity, due to the fact that hydrogen is removed from the reaction mixture. Hence, these undesired hydrogen-consuming reactions are suppressed. As a consequence, at 2.5 bar reaction pressure and a normalized sweep gas flow of 50 there is still a high styrene selectivity of 95.2% maintained at 88.8% conversion, which corresponds to a net increase of styrene yield from 57.6% (fixed bed, 1 bar) to 83.8%. It has to be emphasized that in Figs. 5 and 6 with increasing retentate pressure the pressure differential between retentate and permeate, and hence the driving force of hydrogen removal, also increases. However, this is not the main reason for the observed increase of conversion/yield. This has been confirmed by addi-

Ch. Hermannet al./Journal of Membrane Science 136 (1997) 161-172 tional calculations with retentate and permeate pressures being increased in parallel, for example, for 2.5 bar reaction pressure, zero pressure differential, and a normalized sweep gas flow of 50, ethylbenzene conversion is 84.5% at 92.9% styrene selectivity, thus giving a styrene yield of 78.5%; this is only 5.3% below the yield obtained with 1.5 bar pressure differential (83.8%) but 18.7% above the yield at 1 bar reaction pressure (59.8%).

167

situation is once more different at elevated pressure: 2.5 bar reaction pressure and 10 mbar permeate pressure result in predicted 88.9% ethylbenzene conversion and 94.9% styrene selectivity, which corresponds to a styrene yield of 84.4% compared to 57.6% for the fixed-bed at 1 bar. These results are consistent with those obtained for the previous case of inert sweep gas: increasing the reaction pressure first of all accelerates the reaction rate, and this reduces the limiting effect of the reaction kinetics.

3.2. Evacuation of the permeate gas Figs. 7 and 8 illustrate the situation during evacuation of the permeate gas. First, the permeate pressure has to be decreased to less than 100 mbar in order to withdraw hydrogen from the retentate. Again, at normal reaction pressure there is only a small benefit of the membrane, even if the permeate pressure is kept as low as 1 mbar, that is, only 2% increase of conversion and 1.7% increase of styrene selectivity. The

W

Eth~mter4 corer=ion [ ~ ] U

3.3. Oxidation of the permeated hydrogen with air (reactive sweep gas) For the third case, that is, oxidation of the permeated hydrogen with air, similar results are obtained. In Fig. 9, ethylbenzene conversion and styrene selectivity are plotted versus the retentate pressure. For these calculations a molar ratio of air to ethylbenzene of 10 has been used in order to provide sufficient oxygen for complete conversion of the permeated hydrogen (f --- 1 in Eqs. (2b) and (3)). Hence, the bulk partial pressure of hydrogen in the permeate gas is zero. The permeate pressure is kept at 1 bar. Again, at 1 bar

100

iiiii "11-, I

9O

R P 1 0

80

Fig. 7. Pd-coated porous stainless steel. Case II: Evacuation of the permeate gas. Ethylbenzene conversion versus permeate and retentate pressure (T = 620°C, WHSV = 1.0 h-1, S/O = 12).

.> "~ 70 _m Q~

o0 6O ._o 50 Stynm

10

~

-~,F-

40 I30

. . . .

J" Conversion in membrane reactor • Selectivity in membrane reactor ......... Fixed-bed conversion ....'='"" Fixed-bed selectivity

20 1 Ramit prmmu

Fig. 8. Pd-coated porous stainless steel. Case II: Evacuation of the permeate gas. Styrene selectivity versus permeate and retentate pressure (T = 620°C, WHSV = 1.0 h 1, S/O = 12).

2 3 4 5 Retentate pressure [bar]

6

Fig. 9. Pal-coated porous stainless steel. Case III: Oxidation of the permeated hydrogen. Ethylbenzene conversion and styrene selectivity versus retentate pressure (T = 620°C, Pp = 1 bar, n~ir/nEa =

10, WHSV= 1.0h 1, S/O=12).

Ch. Hermann et al./Journal of Membrane Science 136 (1997) 161-172

168

reaction pressure only a small increase of conversion by 1.8% and of styrene selectivity by 1.7% is predicted, whereas at 2.5 bar 90.5% conversion and 95.8% styrene selectivity are observed, which means 86.7% styrene yield as compared to 56.7% for the fixed-bed at 1 bar. Once more, these results are consistent with the observations for the two previous cases of inert sweep gas and vacuum, respectively.

3.4. Influence of catalyst activity The above results indicate that, under the conditions applied, the crucial factor for the performance of the membrane reactor is on the catalyst side rather than on the membrane side. It appears that the reaction at normal pressure is slow compared to the rate of hydrogen removal through the membrane, so that it limits the rate of the overall process (see also Raich and Foley [28]). This is confirmed by calculations with all reaction rate constants multiplied by a common acceleration factor kac~, thus simulating the effect of a more active catalyst without changing its selectivity properties. The results are shown in Fig. 10 in terms of ethylbenzene conversion and styrene selectivity for the membrane reactor and the fixed-bed, both operated at 1 bar. For the membrane reactor oxidation of the permeated hydrogen is assumed. It is evident that with increasing catalyst activity the difference between the performance of both reactor types becomes more pronounced. At kacc -- 1 only 1.8% increase of con-

version and 1.7% increase of selectivity are predicted whereas at kacc = 4 conversion increase is already 10.2% and selectivity increase is 4.4%, representing an increase of styrene yield from 78.4% in the fixedbed to 92.0% in the membrane reactor.

3.5. Influence of membrane properties To explore the effects of the various membrane layers on conversion and selectivity, additional simulations with varying pore size and Pd-film thickness were carried out. The results indicate that the diffusion of hydrogen through the palladium film is the slowest of the different consecutive hydrogen transport steps; therefore it controls the rate of hydrogen removal. This is illustrated in Fig. 11 in terms of calculated partial pressure profiles of hydrogen across the membrane at the reactor exit. There is no significant drop of hydrogen partial pressure across the gas films on both faces of the membrane nor across the macroporous support. For palladium layers above 10 g m the hydrogen transport resistance begins to affect conversion and selectivity significantly, as shown in Table 3. However, a comparison of the values predicted for 1 tam and 10 g m thickness reveals only 0.7% decrease of conversion and 0.8% decrease of selectivity. Table 4 illustrates the influence of the pore size of the stainless steel support. We find no significant change of conversion and selectivity with decreasing

4000 -

90

---

~

¢

80

~

70



O) 60

--4

c o .~

-

¢s

~

02

ca. 1.8 mm~

!

! J i

,~

20

~-1/atrt~-,~

30 - - ~

20

2= 'I i

2000-

---÷

50

50 ~tm

, ca. 0.2 mm j,

= 40

8

i i

TT im

100

~ 0.4

0.6

0.8

1

2

4

6

8

10

q 5 grn j 10 F m

T I

oS reaction side

stagnant gas film

li

porous support

r / Pd-

stagnant

J T

film

gas film

8

sweep side

kacc [-1 Fig. 10. Pd-coated porous stainless steel. Ethylbenzene conversion a n d styrene selectivity versus c o m m o n reaction rate acceleration factor kate. Case III: Oxidation o f the p e r m e a t e d h y d r o g e n (T = 620°C, P? = 1 bar, nair/nEB = 10, W H S V = 1.0 h - l , S / O = 12).

Fig. 11. P d - c o a t e d p o r o u s stainless steel. H y d r o g e n partial pressure profiles across the m e m b r a n e at the reactor exit. Pd-film thickness as a parameter. Case III: Oxidation o f the p e r m e a t e d hydrogen (T=620°C, PR = 2 b a r , PP = 1 bar, n~/nEB = 10, W H S V = l h 1, S / O = 12).

Ch. Hermann et al./Journal of Membrane Science 136 (1997) 161-172 Table 3 Effect of Pd-film thickness on conversion and selectivity. Case III: Oxidation of the permeated hydrogen (T = 620°C, PR = 2 bar, Pe = 1 bar, n~/nEa = 10, W H S V = 1 h - l , S / O -- 12) Thickness of the Pd-film

Ethylbenzene conversion

169

E ~ m

Styrene selectivity

[rtm]

[%]

[%]

1 5 10 20 50

86.2 85.9 85.5 84.1 80.7

97.2 96.9 96.4 95.4 93.3

Table 4 Effect of support pore size on conversion and selectivity. Case III: Oxidation of the permeated hydrogen (T = 620°C, PR = 2 bar, Pp = 1 bar, nair/nEB = 10, WHSV = 1 h -1, S / O = 12) Pore size of the support [nm]

Ethylbenzene conversion [%]

Styrene selectivity [%]

200 100 50 10

85.47 85.45 85.43 85.21

96.37 96.36 96.34 96.18

Built

,sweep

~

~4trcaH 0

0

Fig. 12. Pd-coated porous stainless steel. Case III: Oxidation of the permeated hydrogen (adiabatic operation). Ethylbenzene conversion versus burnt fraction of permeated hydrogen and normalized sweep gas flow. (T = 620°C, PR = 2 bar, Pp = 1 bar, WHSV = 1 h l, S / O = 12).

W

Btml[f~'l~lJl aem"~mea!

Is~ep ~BH 0 0

pore size from 200 nm to 5 nm, which means that the contribution of the support to the overall hydrogen transport resistance should be small. This is consistent with the fiat partial pressure profiles across the support, as depicted in Fig. 11.

3.6. Adiabatic operation To address the problem of heat management, additional simulations were performed for adiabatic operation which is more likely to occur in a real process situation than isothermal conditions. Heat effects clearly are most pronounced for the permeate oxidation mode. Hence, Fig. 12 shows the predicted ethylbenzene conversion as a function of the oxidized fraction f of the permeated hydrogen, defined in Eq. (2b), and the normalized sweep gas flow, that is, the ratio of air-to-ethylbenzene flow; Fig. 13 illustrates the behavior of styrene selectivity. The reaction pressure is 2 bar; permeate pressure is 1 bar, and the inlet temperature of both feed gas and sweep gas is 620°C. With increasingf conversion increases, caused by the acceleration of hydrogen removal, but also by

Fig. 13. Pd-coated porous stainless steel. Case III: Oxidation of the permeated hydrogen (adiabatic operation) Styrene selectivity versus burnt fraction of permeated hydrogen and normalized sweep gas flow (T = 620°C, PR = 2 bar, Pp = 1 bar, W H S V = 1 h 1, S / O = 12).

the growing heat release due to the combustion of hydrogen which finally overcompensates the heat consumption of the endothermic dehydrogenation. The adiabatic temperature change (cf. Fig. 14) is a function of the sweep gas flow rate as well; hence, the increase of conversion becomes less pronounced with increasing sweep gas flow. Moreover, for f > 0.7 an increase of the sweep gas flow rate beyond the amount required for the oxidation of the permeated hydrogen leads to a decrease of conversion. On the contrary, for f = 0 we essentially have the inert sweep gas mode which means that conversion is increased with increasing sweep gas flow. A comparison of the styrene yield obtained under adiabatic and isothermal conditions is given in Fig. 15. It says that about a 10% higher styrene yield can be achieved in an adiabatic process operated at low sweep gas flow, but this is mainly due to an (unwanted) increase of temperature

Ch. Hermann et al./Journal of Membrane Science 136 (1997) 161-172

170

tw.permtum [K] ~

~

mNl~ a.8 i~mr~¢,~i~q~v~lonf i.I

~ tlownCr/flEB|-I

0 Fig. 14. Pd-coated porous stainless steel. Case III: Oxidation of the permeated hydrogen (adiabatic operation). Reactor outlet tempertature versus burnt fraction of permeated hydrogen and normalized sweep gas flow (T = 620°C, PR = 2 bar, Pp = 1 bar, W H S V = 1 h-', S/O = 12).

Styrene yiek~d~l~nom

sweep idnEB 1-1

m l

0 0 Fig. 15. Pd-coated porous stainless steel. Case III: Oxidation of the permeated hydrogen. Styrene yield difference of the adiabatic

process as compared to isothermal operation versus burnt fraction of permeated hydrogen and normalized sweep gas flow (T = 620°C, PR = 2 bar, Pp = 1 bar, WHSV = 1 h-1, S/O = 12). which at the same time is associated with a reduction of styrene selectivity by about 4% (cf. Fig. 13). Comparing the selectivity at constant level of conversion does not show any significant difference between adiabatic and isothermal operation.

taken from literature. The results indicate that the benefits of the membrane with respect to ethylbenzene conversion and styrene selectivity are only small when operating at conditions close to the commercial styrene process; this is due to the fact that the overall process is limited by catalyst activity. However, if for instance, the reaction pressure is increased, the reaction rate is accelerated and this significantly increases conversion. Contrary to conventional fixed-bed operation, the increase of conversion in the membrane reactor is not accompanied by a significant loss of styrene selectivity, which is explained by the fact that the undesired side reactions (i.e. toluene formation) consume hydrogen and are therefore suppressed upon hydrogen removal. Furthermore, the oxidation of the permeated hydrogen with air seems to be more effective than the application of inert sweep gas or vacuum. At 620°C the model predicts up to a 30% increase of ethylbenzene conversion under isothermal conditions when the retentate pressure is increased from 1 to 2.5 bar. The calculated styrene selectivity of 95,8% is about the same as in the fixed-bed at normal pressure. The simulation results for adiabatic operation demonstrate that the coupling of the endothermic dehydrogenation with the exothermic combustion of hydrogen can lead to excessive reaction temperatures. Hence, to avoid a loss of selectivity (coking) or even damage of the catalyst due to thermal stress, a sufficiently high sweep gas flow is important. Based on the encouraging simulation results an experimental study has been launched recently aiming at a verification of the model predictions; one of the most important aspects of this concerns the long-term stability of the membrane and the catalyst which upon removal of hydrogen is exposed to a substantially different atmosphere and hence might deactivate faster than in normal fixed-bed operation.

5. List of symbols 4. Conclusions and outlook Ethylbenzene dehydrogenation in a composite Pd/porous stainless steel membrane reactor was simulated using a reactor model, which accounts for different mass transport mechanisms prevailing inside the various membrane layers, and a kinetic model

A Bo ci Cp

Co

Cross-sectional area [m 2] Transport parameter defined by Eq. (8) [m 2] Molar concentration of species i [kmol m -3] Molar heat capacity [kJ kmol -a K -1] Hydrogen solubility (e.g. in palladium) [kmol m -3]

Ch. Hermann et al./Journal of Membrane Science 136 (1997) 161-172

dcat

ap D DIj

Catalyst particle diameter [mm] Pore diameter [~tm] Sieverts' law constant, defined by Eq. (12) [kmol m -t s -1 Pa 0.5] Hydrogen diffusivity (e.g. in palladium) [m 2 s -1] Effective Knudsen diffusivity [m 2 s -1]

DO

Binary diffusion coefficient of species i and j [m 2 s -l]

da

Width of the annulus [m] Oxidized fraction of the permeated hydrogen [-] Factor, defined in Eq. (2b) [-] Enthalpy of reaction j [kJ kmo1-1] Molar flux [kmol s -1 m 2] Reaction rate acceleration factor [-] Equilibrium constant [Pa] Molar mass of species i [kg kmol 1] Molar flow-rate [kmol s -t] Partial pressure of species i [Pa] Standard pressure 1.013 [bar] Pressure [Pa] Reaction rate [kmol kgc-alt s -l] Radial coordinate [m], tube diameter [m] Gas constant 8.314 [kJ kmo1-1 K -1] Temperature [K] Flow velocity [m s -x] Mole fraction of species i [-] Spatial coordinate [m]

F

J5 AHR,; Ji kacc

K. Mi

h Pi

Po P rj r,R 9t T bl Yi Z

5.1. Greek symbols Ep ld

Pb 7-

Catalyst bed void fraction [-] Dynamic viscosity [Pa s] Stoichiometric coefficient [-] Catalyst bed density [kgcat m 3] Tortuosity factor [-]

Acknowledgements This paper was presented during the Second International Conference on Catalysis in Membrane Reactors, Moscow, September 24-26, 1996, organized by Prof. V.G. Gryaznov, Russian University of People's Friendship, Moscow. Financial support from Siid-Chemie AG, Mtinchen and the Bavarian Catalysis Research Association FORKAT is gratefully acknowledged.

171

References [1] B.K. Abdalla and S.S.E.H. Elnashaie, Catalytic dehydrogenation of ethylbenzene in membrane reactors, AIChE J., 40 (1994) 2055. [2] B.K. Abdalla, S.S.E.H. Elnashaie, S. Alkhwaiter and S.S. Elshishini, On the kinetics of ethylbenzene catalytic dehydrogenation, Appl. Catal., A: General, 113 (1994) 89. [3] J.G.P. Sheel and C.M. Crowe, Simulation and optimization of an existing ethylbenzene dehydrogenation reactor, Can. J. Chem. Eng., 47 (1969) 183. [4] T.I. Mac Dougall, in HP Impact, Hydrocarbon Processing, April 1996, p. 33. [5] K. Mohan and R. Govind, Analysis of equilibrium shift in isothermal reactors with a permselective wall, AIChE J., 34 (1988) 1493. [6] K. Mohan and R. Govind, Effect of temperature on equilibrium shift in reactors with a permselective wall, Ind. Eng. Chem. Res., 27 (1988) 2064. [7] J.G.A. Bitter, Process and apparatus for the dehydrogenation of organic compounds, U.K. Pat. Appl. 2201159 A, 1988. [8] Y. Liu, A.G. Dixon, Y.H. Ma and W.R. Moser, Permeation of ethylbenzene and hydrogen through untreated and catalytically-treated alumina membranes, Sep. Sci. Technol., 25 (1990) 1511. [9] J.C.S. Wu, T.E. Gerdes, J.L. Pszczolkowski, R.R. Bhave and P.K.T. Liu, Dehydrogenation of ethylbenzene to styrene using commercial ceramic membranes as reactors, Sep. Sci. Technol., 25 (1990) 1489. [10] J.C.S. Wu and P.K.T. Liu, Mathematical analysis on catalytic dehydrogenation of ethylbenzene using ceramic membranes, Ind. Eng. Chem., Res., 31 (1992) 322. [11] Y.L. Becker, A.G. Dixon, W.R. Moser and Y.H. Ma, Modelling of ethylbenzene dehydrogenation in a catalytic membrane reactor, J. Membr. Sci., 77 (1993) 233. [12] G.R. Gallaher Jr., T.E. Gerdes and P.K.T. Liu, Experimental evaluation of dehydrogenations using catalytic membrane processes, Sep. Sci. Technol., 28 (1993) 309. [13] E Tiscareno-Lechuga and G.G. Hill Jr., M.A. Anderson, Experimental studies of the non-oxidative dehydrogenation of ethylbenzene using a membrane reactor, Appl. Catal., A: General, 96 (1993) 33. [ 14] W.S. Yang, J.-C. Wu and L.-W. Lin, Application of membrane reactor for dehydrogenation of ethylbenzene, Catal. Today, 25 (1995) 315. [15] E. Gobina, K. Hou and R. Hughes, Mathematical analysis of ethylbenzene dehydrogenation: Comparison of microporous and dense membrane systems, J. Membr. Sci., 105 (1995) 163. [16] J.N. Armor, Challenges in membrane catalysis, CHEMTECH, Sept. 1992, 557. [17] J. Zaman and A. Chakma, Inorganic membrane reactors, J. Membr. Sci., 92 (1994) 1. [18] E. Kikuchi, Hydrogen-permselective membrane reactors, CATECH, March 1997, 67. [19] J. Shu, B.P.A. Grandjean and S. Kaliaguine, Asymmetric PdAg/stainless steel catalytic membranes for methane steam reforming, Catal. Today, 25 (1995) 327.

172

Ch. Hermann et al./Journal of Membrane Science 136 (1997) 161-172

[20] E. Gobina and R. Hughes, Reaction coupling in catalytic membrane reactors, Chem. Eng. Sci., 51 (1996) 4045. [21] S.S.E.H. Elnashaie, B.K. Abdalla and R. Hughes, Simulation of the industrial fixed bed catalytic reactor for the dehydrogenation of ethylbenzene to styrene: Heterogeneous dusty gas model, Ind. Eng. Chem. Res., 32 (1993) 2357. [22] H.W. Deckman, E.W. Corcoran, J.A. McHenry, J.H. Meldon and V.A. Papavassiliou, Pressure drop membrane reactor equilibrium analysis, Catal. Today, 25 (1995) 357. [23] VDI-Warmeatlas, 6th Edition, VDI-Verlag, Diisseldorf, 1991. [24] H.J. Sloot, C.A. Smolders, W.P.M. van Swaaij and G.E Versteeg, High-temperature membrane reactor for catalytic gas-solid reactions, AIChE J., 38 (1992) 887.

[25] J. Shu, B.EA. Grandjean, A. van Nesten and S. Kaliaguine, Catalytic palladium-based membrane reactors: A review, Can. J. Chem. Eng., 69 (1991) 1036. [26] NAG Fortran Library Manual, Mark 12, Volume 1, 1987. [27] M. Sch~ifer, M. Peric, Skriptum numerische Strtimungsmechanik, University of Edangen-Niirnberg, 1993. [28] B.A. Raich and H.C. Foley, Supra-equilibrium conversions in palladium membrane reactors: Kinetic sensitivity and time dependence, Appl. Catal. A: General, 129 (1995) 167. [29] T.E. Daubert, R.P. Danner, Physical and thermodynamic properties of pure chemicals, Hemisphere Publishing Corp., New York, 1989.