Simulation of ethylene epoxidation in a multitubular transport reactor

0009~2509/86 S3.00 + 0.00 Pergamm fins Ltd.

SIMULATION OF ETHYLENE EPOXIDATION MULTITUBULAR TRANSPORT REACTOR DAE WON

PARK and GEORGES

Institut de PBtrolkochimie et de Synthbse Organique (Receiued

12 February

IN A

GAU’

Industrielle, 13013 Marseille, France 1985)

has been demonstrated that increased yields of ethylene oxide are obtained with a fixed-bed reactoroperating with multiple injection of oxygen and cyclic adsorption-desorption of the desired reaction

Abstract-It

product. In certain conditions, this reactor simulates the operation of a larger two-tube transport reactor; the latter reactor may, in turn, be used to simulate large-scale commercial multitubular transport reactors. The small fixed-bed reactors may also be used for optimization of the process. The reaction of ethylene epoxidation on a silver catalyst is a good illustration of the advantages of a multitubular transport reactor and of this new method of simulation and optimization.

INTRODUCTION

transport reactor (Yousfi et al., 1973) is a dense phase circulating fluidized bed in which the fluidized gas-solid emulsion is circulating inside small tubes (risers). Its main characteristic is that the catalyst The multitubular

is

systematically

reactivation

moved

or deactivation

to

different

zones

where

of the catalyst and adsorp-

tion or desorption of reaction products may be carried out. If desired, a high ratio of catalyst to gas circulation, up to 200 kg per m3 of gas, may be obtained and the catalyst may undergo a full cycle in less than 20 s. In this sense, the multitubular transport reactor is a cyclic reactor. Due. to the systematic catalyst circulation, combined with the modular structure of the reactor, new operating conditions may be obtained:

secutive reactions). Selectivity to ethylene oxide is very sensitive to the operating conditions, even at relatively low conversion. As will be shown, the main features of ethylene epoxidation in a transport reactor may be simulated by the use of a small tied-bed reactor operating in the chemical reaction controlled regime. THE TRANSPORT

REACTOR

The multitubular transport reactor is essentially a fluidized reactor operating in the fast fluidization regime. As indicated in Fig. 1, the catalyst circulates inside the tubes of a she&and-tube heat exchanger: one of two tubes operates as a riser while the other operates

(4 adsorption (W (4 W W

or desorption of some reaction products; operation in the stoichiometric redox mode, in which a metal oxide reacts first with a hydrocarbon and then is reoxidized in a separate operation; multiple injections of reactants along the riser height; operation in the chemical reaction controlled regime, with no heat transfer and no mass-transfer limitation; use of highly dispersed and active catalysts operating at low temperature under conditions where the catalyst is not self-cleaning.

Some of these modes of operation may be useful in improving the reaction yield, depending on the mcchanism and the scheme of the reaction system. lnstead of postulating a set of reaction mechanisms with their related steady and unsteady-state kinetic equations, it is, of course, much better to illustrate the reactor/reaction interaction with the help of a real reaction system. Ethylene epoxidation on a supported fluidizable catalyst has been chosen as a typical oxidation reaction with a triangular reaction scheme (parallel and con-

+To whom correspondence should be addressed. t43

Fig. 1. Configuration of the multitubular transport reactor. 1, Riser tube; 2, downcomer tube; 3. cooling fluid; 4, Ruidization chamber; 5, perforated plate; 6, reactant inlet; 7, fluidization gas; 8, products.

144

D. W. PARK and G. GAU

as a downcomer. Even distribution of catalyst to each riser is ensured by means of good fluidization of the chamber located at the bottom, and injection of equal flows of the reactant mixture into each tube. Since all the risers are identical, the reactor is, in a certain sense, as modular as the shell and tube reactor used with fixed-bed catalysts. The multitubular fixed-bed reactor may be simulated by a single tube (monotube reactor). Similarly, the multitubular transport reactor may be simulated by two tubes, a riser and a downcomer connected by a fluidization chamber and a disengagement chamber (Fig. 2). The diameter of the transport reactor tubes is larger than that of the fixed bed reactor: 0.1-0.2 m in the first case, 0.025 m in the second. The height may also be greater, ranging from 10 to 20 m. But as previously shown (De Lass and Gau, 1973), the transport reactor always operates in the chemical reaction controlled regime; it is practically isothermal at high catalyst flow rate, due to the thermal inertia of the solid. Consequently, any value of tube height and diameter may be chosen, as long as the gas and catalyst residence times are the same as those in the full size reactor. Simulation of ethylene epoxidation has been undertaken in a small two-tube reactor (Gau, 1981; Ghazali, 1980). The tubes, of height 2.7 m, were of 0.022 m inside diameter for the riser and 0.035 m for the downcomer. The gas velocity was varied from 0.7 to 3 m/s and the catalyst hold-up, i.e. the fraction of solid in the riser, was varied from 4.5 to 22%. With this equipment, the catalyst may be reactivated by oxida-

tion in the fluidization chamber, and steamstripped in the downcomer to desorb ethylene oxide from the support. The operation of this reactor is not as easy as it appears since a relatively large amount of catalyst, of the order of 3 1, is necessary, and irreversible deactivation or poisoning of the catalyst may occur frequently during an experimental programme. Fortunately, as will be shown, simulation with a small fixed-bed reactor operating in the chemical reaction controlled regime is also possible. Model of the transport reactor First, we will compare the equations of the transport and fixed-bed reactors in the case where there is no injection of reactants along the tube and no adsorption of reaction products on the solid. Neglecting the acceleration zone of the solid at the entrance of the tube, the fully developed flow corresponds to a zone where a (fraction of solid), us (gas velocity) and II, (solid velocity) are independent of height. This assumption is valid if the pressure drop in the riser is small compared to the absolute pressure, and if the volume variation associated with the stoichiometry of the reaction is small. For the differential element of height dZ, the following equation may be written: (1 -a)u,dc,

= -apsrAdZ

(1)

where ps is the density of the catalyst, C~ is the concentration of reactant A, and rA is the reaction rate of component A per unit weight of catalyst. This equation is very similar to that of a fixed-bed reactor

of solid fraction

B and with differential

height

dZ’: (1 -/_?)u;dc, The two equations fldZ’ (1 -fl)~;

= -/lp,r,dZ’.

are equivalent adz = (1 -a)us

(2)

iE =de

i.e. for equal contact time in the differential reactors. This means that it is possible to simulate any transport reactor of diameter D, height Z and fraction of solid a with a fixed-bed reactor of arbitrary diameter as long as the total contact time (0) is the same in the two reactors:

PZ’ (1 -/?)u;

Fig. 2. Two-tube transport reactor. 1, Riser tube; 2, reactant

inlet; 3, fluidization gas; 4, disengagement chamber; 5, titers; 6, downcomer tube; 7, stripping gas; 8, shell for cooling fluid; 9, cooling fluid; 10, products.

aZ = (1 -a)us

= ‘.

(4)

In general, the best choice for the fixed-bed diameter is 4 mm, in order to reduce radial temperature gradients. In our case, a total length of 2.1 m was chosen: fl being of the order of 0.6, & is calculated from eq. (4). This formula remains valid for small variations of a as a function of Z due to either pressure drop in the riser or volume variation related to the reaction stoichiometry. In this case, the total pressure drop in the fixed bed is chosen to be equal to that of the transport reactor. Simulation with the small fixed-bed reactor, a coiled copper tubing of 2.1 m length and 4 mm inside diam-

Ethylene epoxidation in a multitubulartransportreactor

eter, placed in a constant temperature silicone oil bath, has many advantages compared to the two-tube reactor. The amount of catalyst required is only 25 cm”; operation in the explosive range is safe, and oxidation with air may be simulated, at lower pressure, by oxidation with pure oxygen. Finally, operation at high pressure is also possible. The main disadvantage of this reactor is that the catalyst normally used in a transport reactor, i.e. a fluid&able catalyst of mean particle diameter 60 pm, cannot be used in a fixed-bed reactor because of the high pressure drop and bed segregation (Edwards and Richardson, 1968). Since the catalyst particles operate in the chemical reaction controlled regime, the same support material with a higher mean particle size may be used to prepare the fixed-bed catalyst. In our case, silver was deposited on a macro-microporous silica gel support, of mean diameter 150 pm. A preliminary study has shown that the diameter does not change the activity and selectivity of the catalyst since the silver crystallites are deposited homogeneously inside the porous structure. Optimization of the reactant injection prqfile Since a multitubular transport reactor has only a relatively small number of tubes (less than 200 risers) and no packed bed inside them, it is possible to inject air, or oxygen, at many levels in the tubes, using a small feed tube with calibrated holes. It will be shown below that a high oxygen partial pressure is detrimental to the epoxidation reaction and favours total combustion. A similar observation was made during a study of benzene oxidation to maleic anhydride (Range1 Cordova and Gau, 1983). Furthermore, in some cases, gradual injection of air is the only means of avoiding operation in the explosive range. The fixed-bed reactor used for the simulation may also be used to simulate multiple injection operation. As an example, oxygen was injected at three equidistant points in a 2.1 m reactor: the first point, with flow rate F,, was at the entrance of the reactor where ethylene is also introduced (Fig. 3). From the previous analogy, each length is equivalent to a transport reactor of the same contact time; the equivalent length of the transport reactor can be determined if a hydrodynamic correlation relating solid hold-up (fraction of solid) to gas velocity is known. Such a correlation has been established by Yousfi and Gau (1974). The experimental arrangement may be generalized to a larger number of injection points and simulation of the continuous injection profile is thus obtained.

Q*

5

01 f2

0.

3

Fig. 3. Schematic diagram of multiple injection of oxygen in a tied-bed reactor. F~is the ethylene flow rate. F1, F2. F3 are the oxygen flow rates at points 1, 2 and 3.

145

Optimization of the profile may be undertaken on the fixed bed and, using the hydrodynamic correlation, the optimized profile for the transport reactor may be obtained.

Simulation of ethylene oxide adsorption-desorption As will be shown later, ethylene oxide adsorption on a porous support may be used to increase selectivity. Equations (1) and (2) are not valid if one of the products or reactants is adsorbed in the differential element of height dZ. But, in the limiting case, where most of the ethylene oxide is adsorbed (large adsorption equilibrium constant and/or large quantity of adsorbent), the rates of reactions 1 and 2 are independent of the small ethylene oxide concentration in the gas phase. The transport reactor and the fixed-bed reactor behave in the same manner, as previously described, at least during a sufficiently short pulse, during which the ethylene oxide concentration of the gas is negligible in all points of the fixed bed. The amount of ethylene oxide produced during the pulse is then obtained by desorption (desorption pulse). The solid used for adsorption may be either the support of the catalyst or a mixture of any granular solid of large specific internal area with the catalyst. The fixed-bed reactor is filled with the catalyst + adsorbent mixture; it is operated in the cyclic mode with a pulse for reaction + adsorption and a pulse for desorption.

EPOXIDATION

ON

A HIGHLY

DLSPERSED

SILVER

CATALYST

Ethylene epoxidation with air or oxygen was selected to illustrate the method of simulation of the transport reactor and to test the sensitivity to reactant injection profile, product adsorption and catalyst regeneration. The triangular reaction scheme is as follows:

z)Irz

,I1

r&z

CO1 + H,O where rlr rz and r3 are the corresponding reaction rates. In a typical industrial fixed-bed reactor operating at 20 bars and with pure oxygen, the partial pressures at the inlet are P, = 6 Bars, PO = 2.4 Bars, an inert gas (methane) being used to avoid operation in the explosive range. In order to reduce the loss of ethylene oxide by the consecutive reaction, conversion of ethylene per pass is maintained at less than 15 %, and typical relative reaction rates are: ri = 80, r2 = 20, r3 = 3. For this simple reaction system selectivity may be defined as (Froment and Bischoff, 1979; Satterfield, 1980):

146

D. W. PARK

s= zz

ethylene oxide produced ethylene consumed r1 -r3 = 0.77. rl + r2 + r3

Even at low conversion the consecutive reaction makes a significant contribution. At very low conversion (rj =Z 0) the selectivity is equal to:

and depends only on the ratio rl/r2 (selectivity ratio). The highest selectivity obtainable in practical situations is related to the reaction mechanism (Kilty and Sachtler, 1974): 6C2H., + 60,

+ 6C,H40

lCZHL + 60,

+ 2C0,

7&H,

+ 60,

+ 2H20

+ 602 + 6C2H,0

+ 2C02 + 2H;0

O,, represents atomic oxygen formed on thecatalyst by the reaction of ethylene and molecular oxygen. Therefore, the limiting value is s = 6/7 = 0.857 (85.7 %)_ Higher selectivity cannot be obtained unless oxygen atoms recombine to form the molecular species. The best methods of promoting the recombination of oxygen atoms seem to be operation at low temperature (surface migration of atoms has a low activation energy) or the ‘use of a stoichiometric redox process. This latter method is a cyclic operation: 6Ag

. . .02+6C2H4d6Ag...

6Ag . -.o

$30,

+6Ag.-.02

0 + 6CzH,0

and G. GAU It was hoped that a low-temperature process would either favour some recombination of oxygen atoms, or bring the ratio ri/(rr + r2) to a value close to 6/7, or even reduce the ratio r3/(r1 + rZ), i.e. suppress the ethylene oxide combustion reaction. But the low temperature is necessarily associated with a very high silver specific area. The commercial catalysts used for fixed-bed operation are composed of an inert support of low internal spcciiic area (macroporous), with a silver content of 10% by weight and silver crystallites of 3OOOA. They are of low activity, and consequently high partial pressures and high temperature (260°C) must be used. One of the main advantages of fluid&d reactors is that the catalyst is of a very small size (diameter 60 m compared to 3 mm in the fixed-bed reactor). To increase the silver content of the catalyst (up to 40 y0 by weight) and reduce the size of the crystallites (down to 3OOA), a large internal specilIc area of the support is necessary otherwise crystallite sintering is unavoidable. The support should have not only macropores but also micropores. Mass-transfer resistance with the large particles of a fixed bed would appear even for the same rate per unit mass of catalyst. In fact, as the Thiele modulus indicates: d_

Ik..

for identical values of k, and D,, the modulus would be increased by a factor of 50 if a fluid&d-bed catalyst particle (dp = 60 pm) was replaced by a fixed-bed catalyst particle (d, = 3 mm).

(reduction) (oxidation)

in which the reduction of the surface oxide and its oxidation occur in two separate pulses. In the oxidation pulse, and in the absence of ethylene, the oxygen atoms have a better chance of recombining. The stoichiometric redox process has already been proposed for olefin ammoxidation with mixed oxides in a process known as the “oxidant process” (Callahan et al., 1970). It has also been proposed for olefin oxidation to ketones on tin-molybdenum oxides (Stamicarbon, 1974). An attempt has been made to produce ethylene oxide and propylene oxide by this method (Park et al., 1983; Balzhinimaev et al., 1984) and a 0.94 selectivity in ethylene oxide has been obtained. But, on silver, the coverage with molecular oxygen at the end of an oxidant pulse is always very low, this species being too labile (Park et al., 1983). Even with a silver catalyst of large specific area the oxygen available per pass (i.e. per cycle) is much too small. Instead of using the metal oxide as a reactant it was decided to use silver in the normal catalytic mode, i.e. in the presence of both ethylene and oxygen: the same large specific area catalysts were studied in their application to low-temperature catalytic processes.

The method of preparation of a 40% weight silver catalyst with crystallites of 3OOA deposited on a high specific surface area (60 mZ/g) of silica gel has been described (Ghazali et al., 1983; Gau, 1982). The rate of reaction at 180°C and low partial pressures is about the same as that of the commercial catalyst operating at 260°C and high partial pressures. This catalyst is self-cleaning at high temperature and high oxygen partial pressure, but imder these conditions it has a low selectivity. At low temperature and low oxygen partial pressure, the selectivity is good but coverage of the silver surface with reaction intermediates (deposits) of the total combustion reaction (Park and Gau, 1985) causes strong inhibition. Consequently the catalyst is not self-cleaning. It has been observed to be completely covered by deposits in the case of propylene epoxidation (Park et al., 1983). Operation in a transport reactor or in a cyclic fixed-bed reactor could be used to reactivate the catalyst periodically by combustion of the deposits. The effect of this reactivation has been shown in previous publications (Park et al., 1983; Balzhinimaev et al., 1984; Park and Gau, 1985). A kinetic study (Park and Gau, 1985) has been made

Ethyleneepoxidation in a multitubular transport reactor in order to define the rate equations of reactions 1 and 2 at low conversion. By use of an automated squarepulse cyclic reactor the unsteady-state rate equations were also obtained. They take into account the deposit build-up on the catalyst as a function of time. But even for such a simple reaction scheme it has been shown that at high conversion, no reliable rate equation can be obtained. Consequently, it is necessary to rely on direct experimental simulation using micro-pilot arrangements, such as the small fixed-bed reactor described previously.

SIMULATION

WITH

THE

INTEGRAL

FIXED-BED

REAmOR

The operation of the fixed-bed reactor was computercontrolled as indicated with the semidifferential reactor (Park and Gau, 1985). The simulation of the transport reactor with a fixedbed reactor is valid only if the latter is operating in the chemical reaction controlled regime. Operation in this regime can be obtained by the combination of a lengthy tube (2.1 m) with a small diameter (4 mm) and a small catalyst particle size (150 pm), as indicated by the calculation of Mear’s criteria (Mears, 1971). The catalyst was diluted with quartz particles (10 g quartz per 1 g catalyst). The effects of the various gradients corresponding to Mear’s criteria-radial thermal gradient in the reactor, thermal gradient around the catalyst particle and thermal gradient at the tube wall-have been compared. In our case, it appeared from this comparison that the strictest thermal criterion is the radial thermal gradient:

When the two sides are equal, the rate of reaction is 5 y0 higher than that for the isothermal case. Some examples of the results obtained with the integral fixed-bed reactor are given in Table 1. In the most exothermic case, run No. 4, the two terms of the criterion were 4.4 x low4 < 1.2 x lo-*, therefore the reactor may he considered isothermal. For this run, the Weisz and Prater (Weisz and Prater, 1954) criterion

(rJobsd2, 36 D,C”

147

is also satisfied and consequently there is no masstransfer limitation. The runs of Table 1 were obtained with a catalyst partially covered with chlorine: 1,2_dichloroethane was used as an inhibitor and graduated amounts of this product were injected into the reactant mixture by means of a diffusion tube (Ghazali et al., 1983). Once the surface is sufficiently covered, the inhibiting effect remains for a long time. The ratio of ethylene and oxygen partial pressures was varied as well as the total flow rate to the reactor. There was a marked decrease of selectivity at higher conversion and this result shows that selectivity is essentially controlled by the consecutive reaction. The highest ethylene conversion was 19% (run 1); but it is obvious from the values of the criteria that chemical controlled conditions could be maintained at even higher conversion. Higher conversion may be obtained by reducing the dilution with quartz or the reactant flow rate. Steady-state

integral reactor with multiple injection

If air is used as the reactant instead of pure oxygen it is possible, by use of multiple injection, to maintain a high ethylene to oxygen ratio at the entrance of the reactor and a low ratio at the outlet. In this way, operation in the explosive range may be avoided at all points of the reactor. This can be simulated using the experimental arrangement of Fig. 3; the data obtained in two runs are given in Table 2. An unchlorinated catalyst was chosen because the selectivity variations are amplified without the inhibitor. Ethylene was introduced at the entrance at a flow rate of 0.35 l/h. The total flow rate of oxygen was 0.27 l/h; the oxygen flow rate at the entrance was smaller in the first run and larger in the second. Selectivity was 5.4 points higher in the first run while ethylene conversion was practically the same in both cases. Therefore, it is obvious that the fixed-bed reactor may be used not only to simulate the response to an arbitrary profile but also to optimize the selectivity of ethylene conversion to ethylene oxide. Unsteady-state

cyclic

operation

with ethylene

A reduction of the contribution of the consecutive

Table 1. Influence of the conversion and partial pressure on the steady-state reaction rates (data were obtained with a chlorine-inhibited catalyst at 160°C) Run

1 2 3 4 5 6 7

0.5 0.5 0.5 0.5 0.8 0.8 0.8

(Lz, 0.5 0.5 0.5 0.5 0.2

oxide

adsorption

cz2x10-6a1

(ZE)

of

oxygen

(2) 0.5 0.75 1.0 ;:L: 1.0 2.5

(r, - rJ) x lo5 (molk h)

(r2 + rs) x 10’ (molk h)

1.44 2.06 2.32 3.23 0.58 0.8 1 1.27

0.46 0.58 0.61 0.69 0.16 0.18 0.24

P,

x 102 (atm)

(L

7.12 6.79 5.75 2.67 2.88 2.01 1.25

75.9 77.9 79.3 82.3 78.2 81.8 83.9

D. W. PARK and G. GAU

148

Table 2. Steady-state operation with multiple injection of oxygen (data wereobtained with an uninhibited

catalystat 160°C)

Oxygen flow rates (l/h) F1

Fz

FJ

0.06 0.12

0.09

0.12 0.06

0.09

Ethylene flow rate, FE

(l/h)

0.35 0.35

Conversion Ethylene (%)

Oxygen (%I

24.2 23.7

63.1 62.9

reaction may be obtained by adsorbing ethylene oxide on a solid surface as soon as it is produced. Ethylene oxide may be adsorbed either on the catalyst support or on any granular solid of large specific surface area mixed with the catalyst particles. A silica gel (Spherosil XOA 400, obtained from Prolabo, France) of 400 m2/g was used as the adsorbent, while a silica gel of the same type (Spherosil XOC 30,60 m2/g) was used as the catalyst support. A catalyst prepared with the first silica gel could have been used both for reaction and adsorption. At temperatures lower than 200°C there is very little decomposition of the adsorbed ethylene oxide; the lower the temperature, the higher the adsorption and the lower the decomposition. This is another advantage of a low-temperature process. An adsorption experiment with the automated square-pulse cyclic reactor (Park et al., 1983; Park and Gau, 1983) is presented first. This run was not carried out at high conversion and is only presented for illustration. Chlorinated catalyst (1 g) was mixed with 1 g of silica gel (XOA 400) adsorbent. The reactor was submitted to a two-pulse cycle as indicated in Fig. 4. During the first pulse (R pulse) of 60 s duration, a mixture of reactant gas (50% CzH4, 50% 0,) was injected into the reactor at a flow rate of 0.5 l/h_ This pulse was followed by a nitrogen pulse (N pulse) of 180 s with the same flow rate, and then the cycle was initiated again. Using the method described, the response shown in Fig. 4 was obtained. The ethylene response was a square output pulse with a 30 s delay due to the dead volume of the reactor; CO2 was not adsorbed on the silica gel but the ethylene oxide response was abnormally delayed. This indicates that ethylene oxide is adsorbed and then desorbed during the nitrogen pulse. During desorption, in the absence of oxygen, no decomposition product was formed and consequently, at this temperature, ethylene oxide is not decomposed by silica gel. Figure 4 indicates that the equilibrium constant of the adsorption of ethylene oxide is sufficiently high, and that the simulation of the transport reactor by a cycling fixed-bed is valid essentially for short reactant pulses. Instead of obtaining the response to a cycle by integration of the concentration responses, it is easier to obtain an average composition by analysis of the gas produced in a full cycle. A 55 cm3 container was filled with hydrogen; the flow of gases leaving the reactor was diverted to the

(r, - rj) x lo4 Ime& h)

(rz+rs)x 104 @01/g h)

1.16 0.96

1.68 1.75

40.8 35.4

Fig. 4. Cyclic response curve, with adsorption of ethylene oxide, to a R-N cycle.R (reactant)pulse:50 % CsH4, 50 % 02, r = 60 s; N (nitrogen)pulse.:pure nitrogen, t = 180 s; temperature 170°C.

container during exactly one cycle. The total pressure in the container and the reactor increased only slightly during the reactant pulse. The contents of the container were thoroughly mixed with a magnetic stirrer, and then injected into the chromatograph. A series of experiments was carried out with a fixedbed reactor of 4 mm inside diameter and 0.7 m length, filled with 4.1 g catalyst mixed with 1.5 g silica gel adsorbent. The main factors affecting the selectivity are the flow rate of nitrogen during the nitrogen pulse, and the relative duration of the reactant and nitrogen pulses, as indicated in Table 3. Run 1 shows that cyclic operation with adsorption increases the selectivity to ethylene oxide; this selectivity is 6 points higher than that for steady-state operation (run 5). But from the comparison of ethylene conversion it is also obvious that the catalyst is more active (2.3 times more active) in the adsorption run. This may be interpreted as a reduction of the deposits (reaction intermediates of the total combustion reaction) due to the reduction of r3. The productivity of the reactor during a full cycle is not increased by the same factor. Consequently, the main advantage is the increase of selectivity. If either the duration (run 2) or the flow rate (run 3) of the nitrogen pulse is too small, the ethylene oxide is not efficiently desorbed and not only is the advantage of adsorption partially lost, but a new phenomenon causes a decrease of selectivity. Since an increase of conversion is also observed in these two runs the interpretation is straightforward: the decrease of selectivity is due to the increase of activity associated with lower coverage by deposits.

EthyIcnecpexidation in a multitubulartransportreactor

149

Table 3. Cyclicoperationwithadsorptionof ethyleneoxide at 170°C: iaflueneeof thedurationof the nitrogenpulse and nitrogen flow rate (the composition of the reactantpulse was C2H4 = 52%, OZ = 48%) Run

1 2 3 4 5

Flow rate (l/h)

Duration of pulse (s)

R-pulse

N-pUlSe

R-pulse

0.50 0.50 0.50 0.50

1.90 1.90 0.48 2.25

30 30 30 30

Conversion (% C&L)

(rl - rl) x 104 (mol/g h)

(r* + r3) x 104 (mob h)

22.2 18.3 11.5 19.1

4.44 3.10 1.92 3.66

1.87 2.10 1.35 1.76

70.4 59.6 58.7 67.5

9.7

1.58

0.84

64.4

A)

N-pulse 20 10 20 15

Steady state, with a flow rate of 0.50 l/h (C+H. = 52%, 02 = 48%)

Cyclic experiments (Park and Gau, 1985) have shown that after an ethylene oxide pulse an increase of selectivity is obtained. Consequently, there are two opposite effects:

(1) the adsorption reduces the rate of combustion of ethylene oxide; (2) the deposits produced by this reaction are also reduced and, at lower coverage, the catalyst is less selective. There seems to be an optimum corresponding to the experimental conditions of run 1 or run 4. A full optimization has not been developed. Instead of inhibiting the catalyst with the products of the ethylene oxide oxidation, it would probably be much better to adsorb and desorb ethylene oxide completely and to inhibit the catalyst by chlorine or CO2 coverage. CONCLl_JSION The multitubular transport reactor operates in the chemical reaction controlled regime and therefore it may be simulated using a reduced scale two-tube reactor with the same gas and catalyst residence times. A somewhat cruder simulation may be obtained with a small fixed-bed reactor operating either at steady state or with concentration cycling. The use of this reactor has shown that multiple injection of oxygen and ethylene oxide adsorption on silica gel may be used to increase the selectivity to ethylene oxide. The two reactors may be used not only for simulation but also for optimization. A direct search optimization of the process combining multiple injection, product adsorption, catalyst cleaning and catalyst inhibition could be undertaken. It is expected that an air oxidation process, more efficient than the oxygen process actually used for ethylene epoxidation, may be developed by this method. Other reaction systems may in turn be more sensitive to this optimization than ethylene epoxidation: a study of these reactions in a cyclic fixed-bed reactor with multiple injection is certainly justified. Another characteristic of a multitubular transport reactor process is that it may be based on a large internal specilic area catalyst, with highly dispersed metals or oxides and consequently operated at low temperature. At low temperature, with efficient cata-

lyst circulation, processes of the stoichiometric redox type could be developed. Alternatively, catalysts that are not selfcleaning at low temperature but which may be regenerated by oxidation may be used. The fluidized-bed reactor may also be used of course for low-temperature processes but the systematic circulation of the catalyst in different zones is more difficult to obtain than with the multitubular transport reactor. It can be concluded finally that the reaction of ethylene oxidation on a silver catalyst is a good illustration of the various advantages of this new reactor. NOTATION molar concentration of species A, mol/l c* D diameter of tubular reactor, cm effective diffusivity in a porous catalyst, D, cm*/s particle diameter, cm d, E activation energy, cal/mol F,, F2, F, flow rate of oxygen at points 1.2 and 3.1/h flow rate of ethylene, l/h heat of reaction, cal/mol effective thermal conductivity, cal/s cm “C partial pressure of component i, kPa radius of tubular reactor, cm reaction rate per unit bed volume, mol/s cm3 reaction rate of component A per unit mass, mol/s g reaction rate (ethylene to ethylene oxide), mol C,H,/s g reaction rate (ethylene to CO2 and H,O), mol C2Hl/s g reaction rate (ethylene oxide to CO2 and H,O), mol C2H40/s g observed reaction rate per unit volume of catalyst, mol/s cm3 selectivity, o/0 absolute temperature at the wall of the tubular reactor, K linear velocity of gas, cm/s linear velocity of solid, cm/s height or length of reactor, cm Greek o!

letters

fraction of solid in a transport reactor

150

B PS

D. W.

PARK and G. GAU

fraction of solid in a fixed-bed reactor density of solid, g/cm3 contact time, s Thiele modulus

REFERENCES

Balzhinimaev, B. S.. Park, D.

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