Simulation and optimal design for the residual oil hydrodemetallation in a cocurrent moving-bed reactor

Pergnmon

Chmicnl Eylinee~ing Scimce, Vol. 49, No. 8, pp. 1175-1183. 1994 Copyright 0 1994 Elwicr S&m Ltd Printed in Great B&in. All rights mewed m%2549p4 s6.00 + 0.0

SIMULATION AND OPTIMAL DESIGN FOR THE RESIDUAL OIL HYDRODEMETALLATION IN A COCURRENT MOVING-BED REACTOR CHIEN-LIH CHIANG and ZHONG-REN FANG Department of Chemical Engineering, National Cheng-Kung University, Tainan, Taiwan, 70101, R.O.C. (First

received

4 January

1993;

accepted

in

revisedform 16 September

1993)

Abstract-A restricted diffusion model was set up to investigate the simulation and optimal design of the residual oil hydrodemetallation in a cocurrent moving-bed reactor. The system parameters examined included the Thiele modulus, oil retention time, flow rate ratio and pore size.Three types of catalyst with diRerentactivityprofiles were investigated, i.e. uniform, linear and two-step. Several figures were presented to predict the outlet metal concentrations and deposit conditions of metal sulfide in the outlet catalysts. Based on minimizing the outlet metal concentration, the optima1 catalyst pore sizes and their corresponding optimal reactor lengths were determined. The results show that the optimal pore size increases with increasing values of the Thiele modulus. The usage of nonuniform activity catalysts and multistage reactor can improve the reactor performance. With the values of oil retention time and flow rate ratio properly chosen, it is found that the one-stage reactor with catalysts under the optimal conditions can significantly improve the reactor performance.

INTRODUCTION

Heavy residual oils are now being upgraded to lighter and more valuable products through hydrotreating processes. During the processing, it is known that asphaltenes in the oils have a detrimental effect on the activity and life of the catalysts and that their presence at large concentrations makes treatment extremely difficult. Most of the metal compounds (mainly vanadium and nickel) are concentrated in asphaltenes, so hydrodemetallation (HDM) of these compounds leaves deposits of mixed metal sulfides in the catalyst pores. These accumulated metal deposits in the catalyst pores will hinder the diffusion of the metal-containing molecules. Catalysts are usually replaced or regenerated when the catalyst pore mouth is completely plugged. Hydrodesulfurization (HDS) is the main pretreatment in hydrotreating the heavy residual oils. During the HDS step, the deposit of metal sulfides in the catalyst pores will not only reduce the catalyst activity, but also shorten its life. To cope with this problem it is well recognized that a multicatalyst system (Neilsen et al., 1981; Schuetze and Hofmann, 1984) serves the purpose most satisfactorily. In such a system, a HDM section preceding the HDS section creates the best option for reducing the catalyst cost. In general, the guard HDM reactors are designed to reduce the amount of metals (V + Ni) by 60-70%. so that the downstream HDS reactors can be protected from metal poisoning. The most common arrangement for HDM uses a trickling fixed-bed reactor. However, for the feedstocks with higher metal contents [from 100 to 1500 ppm, Van Zijll Langhout et al. (1980)], an excessively

large

inventory

of catalyst

would

be re-

quired to achieve an acceptable unit-running time. To improve this, Shell developed the so-called

bunker-flow reactor (Van Ginneken et al., 1975), which was actually a moving-bed reactor. This design only needs a small catalyst inventory, and allows replacement of the catalysts without interrupting the operation of the unit. HDM catalysts are designed to have very high metal-removal activity and storage capacity and are selective towards metal removal, asphaltenes conversion and cracking, and also achieve limited desulfurization. It is realized that a catalyst pellet with smaller pores has a larger surface area, however, due to the metal deposit on the catalyst pore surface, it also has a larger diffusion restriction and a shorter lifetime. Thus, there exists an intermediate pore size which yields the optimal activity. On the other hand, catalyst life and deposit capacity could be improved by employing the catalysts with nonuniform activity profiles. Limbach and Wei (1988) found that a centerloaded activity catalyst, which has a higher level of activity near the catalyst interior, could make the pinch point (a point where complete pore plugging happens first) shift from the pore mouth toward the catalyst interior. Their results indicated that even with a small change in the activity profile, the deposit capacity could be significantly improved. Numerous studies have been made (Newson, 1975; Oyekunle and Hughes, 1987; Rajagopalan and Luss, 1979) to describe the restricted dimusion of the large metal-containing molecules in the catalyst pores and to predict the catalyst performance. In addition to the simulation, Rajagopalan and Luss (1979) investigated the optimal pore sizes for the spherical and slab catalyst pellets. There are, however, only few studies given to the simulation and optimal design of the HDM reactor. Do (1984)examined the optimal catalyst pore size profile along the length of the fixed-bed reactor by considering the initial activity only. Chiang and Tiou

1175

1176

CHIEN-LIHCHIANGand ZHONG-REN FANG

(1992) studied the optimal design for the residual oil hydrodemetallation in a fixed-bed reactor by using the uniform, linear, egg-yolk and edge-reduced activity catalysts. As mentioned above, moving-bed reactor is one of the best choice to solve the problem of which the metal contents in the feedstocks are high. The aim of this study is to place emphasis on the simulation and optimization of the moving-bed reactor in the industrial HDM process. In the present study the restricted diffusion model proposed by Rajagopalan and Luss (1979)was used to describe the metal deposition on the catalyst pores. Then the local demetallation rate is incorporated into the bulk mass balance to establish the whole reactor model. For the uniform activity catalysts, several figures are presented to predict the reactor performance. It is noted that the principles behind these figures are that the catalysts leaving the reactor are to be regenerated or replaced, which is a reasonable design criterion from the economic viewpoint. The catalysts

investigated in this study include the uniform, linear and two-step activity profile catalysts. The optimal pore sizes are obtained by minimizing the reactor outlet metal concentration. In addition to the optimal design in one-stage reactor, the optimal pore sizes and their corresponding optimal lengths in the two-stage reactor are also examined.

By solving eqs (1) and (5), the instantaneous demetallation rate in a single pore of half-length I can be calculated by R,(1, t) = 2nrP(0)lkCb0q(t)

(8)

where

Based on the ideal pore model (Wheeler, 1951), the demetallation rate per unit volume of a catalyst pellet is given by R,,(t) =

E rri(O)l

R,(l, 1).

The total mass of the metal deposit per unit volume of the catalyst pellet can be expressed as w=

s’

sp,

(1 - f’)d<.

0

To explore the effect of nonuniform activity catalyst on the hydrodemetallation performance, two types of nonuniform activity profiles are considered, i.e. linear and two-step. Their mathematical expressions can be written as k(5) = &E)

(12)

h(5) = s(S - 0.5) + 1, 0 6 [G 1, MATHEMATICALMODEL

Most of the demetallation rate is first order with respect to the metal concentration (Rajagopalan and Luss, 1979; Shimura et al., 1986; Limbach and Wei, 1988). Hence, only first-order kinetics is considered. The restricted diffusion behavior in a single pore can be described by the model proposed by Rajagopalan and Luss (1979) as

0 6s 6 2 (linear)

(13)

I

1 h(5) = (

C2U -P) •t PI’ 2

\ cw - PI + PI’

0 6 5 < p (two-step) (14) PC 56

l,O$

p$l.

In order to compare the performance among the various catalysts, the average activity of the nonuniform activity catalyst is made the same as that of the uniform activity catalyst. It is worth pointing out that the two-step profile catalysts have their activities reduced at the outer portion of the catalyst, and they du can be made through competitive adsorption tech-=o, <=I dS niques (Komiyama, 1985). The linear and two-step activity profile are shown in Fig. 1. The activity ratio where the Thiele’s modulus, $, is defined as for the two-step profile is fixed as two for simplifica4’ = 2kP/r,(O)D,(O). (4) tion, and the parameter p is a dimensionless distance The symbol y in eq. (2) represents the dimensionless from the pore mouth. The parameter s in eq. (13) is the slope of the linear profile; its upper limit is set equal to bulk metal concentration outside the catalyst pellet. The rate of the metal deposit per unit length of the two. The cocurrent bunker-flow reactor proposed by pore can be evaluated by Shell (Van Ginneken et al., 1975) consisted of two dj uv sets of sluice systems and rotary valves located at the -= -(5) dt r,(O) top and bottom of the reactor, respectively. This exquisite design enables the addition and withdrawal with initial condition of the catalysts, and rotary valves are used to regulate f=l, t=o (6) the amount of catalysts to be transported. Detailed operating procedures were described in their report. where u is the metallation velocity, i.e. Bunker-flow reactor can be considered as a moving@k&o M bed reactor, in which the steady-state condition is also “(7) satisfied. The metal content in the residual oil in Pm

Residualoil hydrodemetallalionin a oncurrent moving-bed reactor lessvariables:

2.0 -

1.1 1.6 -

1177

(I)LiBeu.S=2 (2) Two-rtep. P = 0.5

y+ LO

1.4 -

xc-

1.2 -

(23)

X L

W=W PC

(24) (25)

eq. (19) can be rewritten as y=l-kw. 0

0.2

0.4

0.6

0.8

1.0

C,,aMr

(26)

Substitution of eq. (7) into eq. (26) yields

I y&kp.W. “Pm’

Fig. 1. Activity profiles of the catalysts.

(27)

The relationship between the real time variable t and the dimensionless reactor length x can be expressed as practice is in the range L%-1500ppm, and hence we may assume that the reaction is isothermal. Excess of hydrogen is added to the reactor to minimize the coking effect. From this the diffusion resistance in the gas and liquid phases may be neglected when compared with the restricted diffusion in the catalyst pore. In addition, plug flow pattern is assumed for simplification in this analysis. The mass balance for the bulk metal concentration

dt = r, dx.

(28)

The pseudo-steady-state equation (1) is solved by central difference method with 300 unequal partitioned nodal points. Modified Euler’s method is used to solve eqs (5), (11)and (27). Due to the metal deposition, the catalyst pore will be completely blocked iff$ &,. For the case of uniform activity catalysts, complete plugging always occurs first at the pore mouth. Thus, if f$ Lo, y and in the residual oil is as follows: W will remain the same elsewhere in the reactor. However, for the nonuniform activity catalyst in U A( 1 - sb)R,(t) (15) which the activity at the catalyst center is larger than that at the pore mouth, complete pore plugging may c, = CbO, x = 0. (1’3 first happen at some points inside the pore. If it occurs On the other hand, mass balance for the loaded metal at 5,) demetallation reaction would continue to proceed in the interval [0,5,] only. In this regard, eqs (3) deposit in the catalysts can be derived as and (9) should be modified as dw Gz = A(1 - eb)Rsl(t)aM (17) (29) w=o, x=0. (18)

$ =-

Combination of eqs (15)and (17) yields

G -& = C&o

-

(30)

C,).

We denote by t, and t,, the retention time of the catalysts and the residual oil in the reactor, respectively, i.e. z, =

Golden section method was used to search the optimal catalyst pore size which yields the minimum outlet metal concentration. In searching for the optimal pore sizes, the initial pore radius should be excluded from the definition of the Thiele modulus. Hence, a modified Thiele’s modulus is defined by

A(1 - && G

(20)

Q’=-.

2lz DABr,

2, r* = r

(31)

RESULTS ANDDLSCUSSION The parameters affecting the reactor performance include the retention time of the residual oil, z,, flow II I=rate ratio. r, feed metal concentration, CbO. Thiele (22) G modulus, @, and initial ratio of molecular to pore where r is the ratio of volumetric flow rate of residual radius, lo. A typical design of the HDM moving-bed oil to catalysts. Introducing the following dimension- reactor aims at a degree of metal removal of about and

CES 49-8-G

CHIEN-LIHCHMNGand ZHONO-REN FANG

1178

50%. The values of the Thiele modulus in practice (Limbach and Wei, 1988) range from 1 to 10; hence three values of Q are considered, i.e. 1,3 and 10. The typical values of the system parameters for the simulations are shown in Fig. 2. The results in this study are divided into two parts. The first part is devoted to the simulation of the HDM moving-bed reactor. Besides elucidating the effect of system parameters on reactor performanffi, some figures are proposed to predict the outlet metal concentration and the condition of metal sulfide deposit in the outlet catalysts. In the second part the results from the simulation are further applied to search for the optimal designs of the reactor, which include the optimal pore sixes and optimal combinations of the reactor arrangement. For the considered case of cocurrent moving-bed arrangement the change of the metal concentration would be the largest at the reactor inlet, then decrease along the length of reactor. These results are expected since the highest metal concentration and fresh catalysts are located at the reactor inlet. With the progress of reactor length, the amount of metal deposit on the catalyst surface increases, which in turn causes the demetallation rate to decrease. In the following discussion, uniform activity catalysts are employed to examine the reactor simulations. Application ofthe theory Figure 2 shows the outlet metal concentration, y,, and effective pore size at the pore mouth of outlet catalysts&, with r,, and r as the parameters. For the case of uniform activity catalysts, it is realized that the largest metal deposit rate is located at the pore mouth; hence we may use the effective pore size at the pore mouth,f,, to represent the deactivation condition of catalysts. r, is the residence time of residual oil in the reactor, while I is the ratio ofoil flow to catalyst flow. The product of t, and r gives the residence time of

catalysts in the reactor. With higher values of T, and lower values of r, Fig. 2 shows that the outlet metal concentration, ye, could be considerably reduced, The reason for this is that the longer residence time of oil (larger r,) and shorter residence time of catalysts (smaller r) would favor the removal of the metal content in residual oil. It is noted from the figure that, at smaller values of r,fI may be greater than 1,. This means that the catalysts arc not completely plugged before they leave the reactor. The outlet catalysts are either replaced or regenerated in the so-called bunkerflow process. From this point of view, the system parameters are so chosen that the catalysts are utilized as efficiently as possible, i.e.fe = Ize.Figure 2 also indicates that, for each pair of ‘I, and An, there exists a value of r which makes the catalysts just being completely blocked at the reactor outlet. In the following simulations, system parameters are determined according to this condition. Another factor which is important in the reactor design is the feed metal concentrations, CbO. It is noted that the concentrations in eqs (1) and (27) are expressed in the dimensionless form. Also it can be found from eq. (7) that the metallation velocity is proportional to kCbO.These indicate that the relatidn between the feed metal concentration and reactor performance would be a simple linear one. These phenomena are indeed observed in our numerical results. For example, if C,, = 0.005 kmol/m3 (about 500 ppm), Q, = 3, &, = 0.1 and T, = 2, it can be found from Fig. 2 that the value of r is equal to 1090, with which the pore mouths of the outlet catalysts are just completely blocked, i.e. fe = L,,,. If Cbo is doubled, catalyst lifetime would be reduced to half of the former case, and its corresponding value of r to meet the same conditions is also halved, i.e. 545. From these results, we may preclude the variable Cao from the simulations. As mentioned above, the design criterion here is that the pore mouths of the outlet catalysts are just completely plugged. Figures 3,4 and 5 plot y, vs 1, for 0.70-

t

YC

Y.

0

I

I

5ocl

KKlo

I

1500

I

2ocm I3

I Fig. 2. y. and /; vs I for uniform agtivitycatalysts (w = 3. &,=0.1,C~o=5x10~Jkmol/ms,k=6x10-1zm/s,r.= 2 x 10e9m, OL = 2, M = 144.5kgjkgmol, p. = 2420kg/ma, p. = 750kg/ma, E= 0.5, b = 0.5).

0.10

I I I 0.15 0.30 0.23

I 0.30

I I 0.3s 0.40

L Fig 3. y, vs r10 for uniform activity catalysts (a = 1, C,, = S x lo-” kmol/m’).

Residual oil hydrodemetallation

1179

in a cocurrent moving-bed reactor

various vales oft, and with Q equal to 1,3 and 10, with Cpequal to 1,3 and 10, respectively. These results respectively, for each figure. Notice that each point in would be further applied to find the optimal designsof the curves of these figures has its corresponding value reactor in the following discussion. Considering the most probable condition in pracofr, with which the pore mouths of the outlet catalysts are just entirely blocked. It can be found through tice, i.e. 0 = 3 and uniform activity catalysts, an atthese figures that there exists a value of &, with which tempt was made to predict their simulation results. its corresponding y, is a minimum, and this do is Figure 6 was obtained by regression analysis from 574 nearly independent of the value of 7,. From the figures numerical calculations, their average errors in predicwe can find that they are equal to 0.3,0.2 and 0.18 tions were below 5%. The application range for A,,

0.20 L

I 0.10

I

I

I

I

I

I

0.15

0.20

0.25

0.30

0.35

0.40

”

._”

Fig. 4. y. vs & for uniform activity catalysts CbO= 5 x 10-s kmol/m3).

(@= 3,

r

\* F 2.0

2.5

‘8.‘t ---- 5

240

‘.

8%

‘c

9.

.5

..

l

-._

--._

---..___

"-0.3

---____:fg

AtJ=o.3

3.0

3.5

o1.0

4.0

1.5

2.0

2.5

3.0

3.5

4.0

70

Fi

720 600 1

.

‘*.

120

'0

fall

0.15

480 360 240 120 0-

‘\

,

LO

’

‘*.

o5

‘. ‘. ‘2, l*

----

%_ -. l._ -.

l.

0.3 -___---. --._

-__:::-fE;:;

i-

1.5

r,

0.4

5,

2.0

0.30

Fig. 5. y. vs do for uniform activity catalysts @ = 10, C,, = 5 x 10m3kmol/m3).

360

1.5

0.20 x0

x,

1.0

\*.__ --._ -rr 0.15

0.10

2.5 3.0 70

r

0.4 &#=o.z 33

4.0

Fig. 6. y. andf. as a function ofr and 7. for the uniform activity catalysts (Q = 3, C,, = 5 x 1Om3kmol/m3)

1180

CHIEN-LIHCHIANG and ZHONG-RENFANG

and ~~in this figure are 0.1-0.4 (200-50 A for r& and l-4 h, respectively. Figure 6 can be used to determine the system parameters to meet a specified reactor performance, which includes the outlet metal concentration and the condition of the outiet catalysts. For example, considering the case of d, = 0.1, the point P in the figure represents an operation result that ye = 0.3 andf, = 0.3 if T#= 2.9 h and r = 750. If other variables are fixed except that the value of r is increased to 940, then the outlet metal concentration willreach 0.4 and its corresponding value of fE is a little smaller than 0.1. This means that the pore mouths of the outlet catalysts are already completely plugged; however, this figure cannot show the location of the reactor in which the catalysts are completely blocked. As described in the Introduction, the aim of this study is to find the optimal pore size to increase the metal sulfide loading. For the nonuniform activity catalysts, besides the optimal pore sires, the optimal activity profiles are also determined, but they should satisfy the activity constraint equations, i.e. eqs (13) and (14). According to the Shell process (Van Ginneken et al., 1975) the requirements in the HDM moving-bed design are that yp is around 0.5 and outlet catalysts are just completely plugged. In this regard, we have to determine r and 2, to satisfy the above requirements for a given pair of @ and &, before searching the optimal pore size. These calculations are performed using the uniform activity catalysts as the basis. For instance, if 1, = 0.1 and @ = 3, it can be found from Fig. 2 that, when r = 1090 and r, = 2 b, y. will be equal to 0.4972 andf, = &. Then we fix r and

rO,and determine the optimal designs for the uniform and nonuniform activity catalyst, i.e. 12, s* or p*, by minimizing the outlet metal concentration ye. Optimal designs in one-stage reactor

lo = 0.1 (rp = 200 A) is a typical value in HDM catalysts. Simulations and optimal designs for this case are listed in Table 1. Results in this table show that the optimal pore sizes for the nonuniform activity catalysts are smaller than those of the uniform activity catalysts. This is attributed to the fact that the interior pore surface of the former can be more effectively utilized. In addition, Table 1 indicates that the introduction of optimal pore sizes and nonuniform activity profiles can improve the reactor performance over the conventional design with uniform activity catalysts, especially at larger values of Q. From Table 1, we may observe that the improvement by the optimal designs are not significant if r, and r are determined according to a typical & of 0.1. To improve this, we may take advantage of the results from Figs 3-5. Satisfying the criteria that y, is around 0.5 andh is nearly equal to Lo, Figs 3-S show that the best combinations of (&, r,,, r) to minimize the outlet metal concentration are (0.3, I.5 572.6). (0.2, 2, 538) and (0.18,5, 219.88) for @J= 1,3 and 10, respectively. Based on these values their optimal designs are shown in Table 2. A comparison between Tables 1 and 2 indicates that, with the proper choice of r, and r, the optimal designs of catalysts, whether uniform or nonuniform activity profile, can improve the performance significantly. In this analysis, the pore radius of the metal-

Table 1. The optimal designs for the catalysts (denoted by an asterisk) in one-stage reactor NO.

r

@=

1, To

=

1.5

0.1 1* 2 2’ 3

3*

1480 1480 1480 1480 1480

0.4834

0.092 0.1 (S= 1) 0.1264(s* = 0.6883) 0.1 (p’ = 0.9) 0.1191(p*= 0.7096)

0.4802 0.4762 0.4563 0.4773 0.4744

d = 3. t. = 2 0.1 0.07847

1* 2 2*

0.1 (S= 2) 0.1123(a* = 1.2832) 0.1 (p = 0.7)

3*

0.10108(p* = 0.495)

0.4972 0.4700 0.4575 0.4309 0.4502 0.4415

@= 10,‘T,= 5 1’

2 2* 3’

420 420 420 420 420 420

0.1 0.056 0.1 (S= 1) 0.0568(a*= 0.0021) 0.1 (p’ = 0.15) 0.069(p’ = 0.1049)

0.5303 0.4459 0.473I 0.4495 0.4660 0.4491

Note: (1) uniform activity catalyst; (2) linear activitycatalyst;(3) two-step activity catalyst.

Residual oil hydrodemetallation

1181

in a cocurrenl moving-bed reactor

Table 2. The optimal designs for the catalysts (denoted by an asterisk)in one-stage reactor

cp= 1, ra = 1.5 1

572.6

1*

572.6 572.6 572.6 512.6 572.6

2 2’ 3 3*

0.2611 0.1536 0.1782 0.1410 0.1719 0.1456

0.3 0.2140 0.3 (s = 2) 0.2406 (SC= 0.9099) 0.3 (p = 0.5) 0.2386 (p* = 0.5849) 0 = 3, r, = 2

1

1’ 2 2* 3 3’

538 538 538 538 538 538

0.2

219.88 219.88 2 19.88 219.88 219.88 219.88

0.18

0.4273 0.3090 0.3416 0.3037 0.3441 0.3057

0.12434 0.2 (s = 1) 0.1354 (s* = 0.6826) 0.2 ( p*

= 0.27) 0.1379 (p’ = 0.2349) 0 = 10, T0= 5

1 1’ 2 2* :*

0.5071 0.3632 0.5178 0.3632 0.4676 0.3633

0.08654 0.18 (s = 2) 0.0863 (s* = O.OODO8) 0.18 (p’ = 0.7) 0.0867 (p’ = 0.9999)

Note: symbols are the same r~sin Table 1.

containing molecules is 20 A. According to the results in Table 2, taking @ = 3 as an example, the optima1 pore diameters for uniform, linear and two-step activity catalysts would be 324,295 and 290& respectively. These values are substantially larger than those reported by Do (19&l), which were determined by considering the initial catalyst activity only. In addition, Shimura et al. (1986) reported an optimal pore diameter of 180A, which was obtained for the vanadium removal by considering the initial catalyst activity. According to the results reported by Rajagopalan and Luss (1979), the optimal pore sizes based on the initial activity were significantly smaller than those based on the total lifetime activity. In this regard, our results obtained here seem reasonable. If the pore radius of catalyst is given, e.g. 0 = 3, i0 = 0.2, Table 2 shows that the introduction of the nonuniform activity profile, either linear or two-step profile, can significantly improve the performance. This phenomena can be clearly observed from Fig. 7, which indicates that the interior parts of nonuniform activity catalysts can be more efficiently utilized. However, if each catalyst is made with its optimal design, Table 2 shows that the benefit of nonuniform activity catalysts over uniform activity catalyst is slight, which can also be illustrated by Fig. 7. The reason for this is that the residence time of catalysts is not long enough to take advantage of the benefit of nonuniform activity catalysts. In addition, it may be noted from Table 2 that s* decreases and p* increases with increasing values of0. For the case of Cp= 10, s* approaches 0 and p* approaches 1, which are near the uniform activity profile. In this regard, the usage of uniform activity catalysts with their optimal pore dia-

0.6 f 0.4

(‘) optimal, ho = 0.2 (1) uniform

0

0.2

0.6

0.4

0.8

1.0

f Fig. 7. /

vs C for the outlet catalysts (@ = 3, C,, = 5 x IO-’ kmol/mg, r. = 2, r = 538).

meters provides a good solution to the design of HDM cocurrent moving-bed reactor. Optimal designs in two-stage reactor In order to compare the performance with onesection design, the total residence time of oil, T,, is kept the same. Also the values of rl, and r1 for the two-section design are determined separately to meet the design criterion, catalysts for uniform

i.e. the pore mouths of outlet activity catalysts are just com-

pletely plugged. To determine the optimal designs of two-section reactor, optimal division points should be considered first. Numerical results show that the optimal division point, x*, is near 0.5, if the same type of

CHIEN-LIH CHIANG and ZHONG-REN FANG Table 3. The optimal designs for the catalysts (denoted by an asterisk)in two-stage reaclor

1

1800 2500

1’

1800 25Ml

2

1800 25oll

2*

1800 2500

3

1800 2500

3*

1800 2500

I

837.8 1324.5

1*

837.8 1324.5

2

837.8 1324.5

2*

837.8 1324.5

3

837.8 1324.5

3’

837.8 1324.5

0.6953

0.1 0.1

0.6773

0.079 0.082

0.4512

0.1 (s = 2) 0.1(5=2)

0.6692 0.4408 0.1127(s* = 1.2845) 0.1172(s* = 1.2622)

0.6501 0.4120 0.6638

0.1 (p = 0.7) 0.1 (p = 0.7)

0.4307 0.1014(p* = 0.4949) 0.1751(p’ = 0.4349)

0.6576 0.4222

Q = 3, T, = 2 0.2 0.2

0.6325 0.4ooo 0.1253 0.1339

0.5482 0.2896 0.5887

0.2 (s = 2) 0.2 (s = 2)

0.3424 0.1365(s* = 0.6352) 0.1394(s* = 0.3472)

0.2 (p = 0.5) 0.2 (p = 0.5)

0.5442 0.2864 0.5843 0.3331

0.1392(p’ 0.1343(p*

= =

0.2249) 0.0085)

0.5460 0.2881

Note: (1) uniform;(2) linear;(3) two-step.

catalysts are used. Also this x* is found to be nearly independent of CD.However x* would be deviated form 0.5 if various kinds of catalyst are adopted in the different sections. But their performance, i.e. outlet metal concentrations, are close to those arranged with the same type of catalysts. Therefore, for simplification, we set x* equal to 0.5 and then search the optimal design for each section. Table 3 shows the optimal designs of two-section arrangement at Q = 3. It is seen from this table that the optimal pore sizes at the upper part of the reactor are slightly larger than those at the lower part of the reactor. This is due to the fact that the former has higher bulk metal concentration; hence larger pore sizes are needed to increase their metal deposit uptake. A comparison among Table 3 and Tables 1 and 2 indicate that the application of two-section design can improve the performance. But these effects are not significant; this is due to the fact that the residence time of catalysts in the moving bed are not long enough to take more metal sulfide deposit. Comparison withfixed-bed design A comparison of optimal designs between the moving-bed and fixed-bed reactors (Chiang and Tiou, 1992) indicate that the optimal pore sizes and total

mass of demetallation of the latter are both larger than those of the former. Furthermore, the nonuniform activity catalysts show significant benefit over uniform activity catalysts in the fixed-bed reactor, even if all the catalysts are under their optimal designs. The major reason for this is the duration of catalysts in the reactor. It is larger than 4OWh in the fixed-bed reactor, while, according to Table 2, it is only around 1000 h in the moving-bed design. In this regard, we may expect that the performance in the moving-bed reactor would be further improved if longer residence time of catalysts are adopted.

CONCLUSION The impact of pore size, reactor arrangement, flow rate ratio, residence time of residual oil and nonuniform activity profile on the performance of HDM cocurrent moving-bed reactor is investigated in this work. Based on the numerical results, several figures are presented to predict the outlet metal concentration and the condition of outlet catalysts. The optimal reactor designs are determined by minimizing the outlet metal concentration, and are subjected to the criterion that the values of r are selected to have the pore mouth of the outlet uniform activity catalysts

Residual oil hydrodemetallation

been just completely plugged. The results presented here indicate that the two-section design and the usage of nonuniform activity catalysts can improve the reactor performance. In addition, improvement of the performance would be significant if the catalysts are designed with their optimal pore sizes.

in a cocurrent moving-bedreactor Greek letters

do E sb i tl

NOTATION cross-sectional concentration

area of the reactor of reactant in the pore,

kg mol/m 3 feed concentration bulk concentration bulk liquid diffusivity, m2/s liquid diffusivity in a pore [ = DAB(l - l)4] initial liquid diffusivity in a pore dimensionless pore size [rjr,(O)] dimensionless pore size at the pore mouth of the outlet uniform activity catalyst dimensionless diffusivity [ = Dp/Dp(0)] volumetric flow rate of the catalysts, m”/h function of activity profile demetallation rate constant, m/s average k for the nonuniform lysts half-length of the pore, m

activity cata-

reactor length, m molecular weight of the metal deposit volumetric flow rate ratio of residue oil to catalyst molecule radius, m pore radius, m initial pore radius, m demetallation rate in a pore defined by eq. (8) demetallation rate per unit volume of pellet defined by eq. (10) time, h dimensionless concentration in a pore ( =

c/c,,)

volume flow rate of the residue oil, m3/h metallation velocity, m/s weight of metal deposit per unit volume of pellet, kg/m3 weight ratio of metal deposit to fresh catalyst dimensionless reactor length distance along the reactor, m dimensionless bulk concentration ( = C,/ CM) outlet bulk concentration distance along the pore, m

1183

number of metal sulfide molecules per molecule of organometallic reactant pellet porosity bed porosity dimensionless distance in the pore ( = z/I) instantaneous effectiveness factor defined by eq. (9) ratio of molecule to pore radius ( = rm/r,) [ = ~.l~,(w optimal &, in ith reactor

catalyst density, kg/m3 metal deposit density, kg/m3 residence time of the residue oil in the reactor, h residence time of the catalysts in the reactor, h Thiele’s modulus defined by eq. (4) modified Thiele’s modulus defined by eq. (31) REFERENCES Chianp. C. L. and Tiou. H. H.. 1992.Ootimal design for the residual oil hydrodemetallation in a fixed bed reactor. Chem. Engng Commun. 117,383-399. Do, D. D., 1984,Catalytic conversion of large molecules: effect of pore size and concentration. 4.I.Ch.E. J. 30, 849-853. Komivama. M.. 1985. Desien and ortoaration of imoreenated catalysts. G&l. Rev. Sci. En& 27, 341-346. I Limbach, K. W. and Wei, J., 1988, Effect of nonuniform activity on hydrodemetallation catalyst. 4.LCh.E. J. 34, 305-313. Newson, E.. 1975, Catalyst deactivation due to wre~lugging by reaction prod&s. lnd. Engng Chem. Pk&D& Dev. 14,27-33. Nielsen. A., Coooar. B. H. and Jacobsen. A. C., 1981. Composite c~talysibeds for hydroprocessi& of h&y r&idua. Div. Pet. Chem. Am. Chem. Sot. Sytnp. Ser. 26,44(.X445. Oyekunle, L. 0. and Hughes, R., 1978, Catalyst deactivation during hydrodemetallation. Ind. EAgng Chem. Res. 26, 1945-1950. Rajagopalan, K. and Luss, D., 1979, Influence of catalyst pore size on demetallation rate. Ind. Engng Chem. Process Des. Dev. 18, 459-465. Schuetze. B. and Hofmann. H.. 1984. How to uosrade heavv feeds. Hydrocarbon Pro& 75-8i. *w r Shimura, M., Shiroto, Y. and Takeuchi, C., 1986. Effect of catalyst pore structure on hydrotreating of heavy oil. Ind. Engng Chem. Fundam. 25, 33&337. Van Ginneken, A. I. 1.. Van Kessel. M. M., Pronk, K. M. A. and Renstrom, G., 1975, Shell process desulfurizcs r&d. The Oil and Gas J. 5943. Van Ziill Langhout, W. C., Ounerkerk, C. and Plonk. K. M. A., 19i0, New process hydmtreats metal-rich feed stocks. The Oil and Gas J. 12C-126. Wheeler, A., 1951, Advances in Catalysis, Vol. 3, p. 250. Academic

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