HDS catalyst regeneration: Coke burning-off by air in a shallow bed

Chemical Erylrnecring Science. Vol. 40. No. 9. pp. 1717-1721. Printed m Great Britain

HDS

KAREL

CATALYST

KLUSACEK,

1985. C

cno!-2509/85 s3.00 + 0.00 1985. Pergamon Press Ltd.

REGENERATION: COKE BURNING-OFF BY AIR IN A SHALLOW BED HELENA

DAVIDOV&

PAVEL

FOTT

and PETR

SCHNEIDER

Institute of Chemical Process Fundamentals, Czechoslovak Academy of Sciences, 16502 Prague 6 Suchdol, Czechoslovakia (Received 4 June 1984) Abstract-Regeneration of coked HDS catalyst by air in a thin layer was studied experimentally and simulated using a simple model of the reactivation reactor. Using kinetic data on deposited coke oxidation obtained separately, the time course of the catalyst petlet temperature was correctly predicted. Conditions under which safe regeneration takes place were determined and verified experimentally as well as conditions

under which coke ignites spontaneously. Under ignition conditions the degree of regeneration can not be controlled, which might result in irreversible catalyst deactivation.

INTRODUCTION

of carbon and sulphur deposits (hereafter denoted as coke) frequently causes deactivation of industrial catalysts. Coke blocks active catalytic sites and catalyst pores; this leads to higher energetic demands of the process, as the decreased catalyst activity has to be compensated, e.g. by an increase of reactor temperature. In reactors the catalyst activity decreases after a certain period so much that it has to be reactivated. During reactivation, the coke is burned off by appropriate reactivating medium, usually by air, diluted sometimes with nitrogen or by steam. The traditional way, based on the reactivation of catalyst inside the reactor does not require manipulation with the catalyst but has, however, several serious disadvantages. The main disadvantage is the difficult temperature control in the catalyst bed during the unsteady-state, strongly exothermic oxidation of coke. Thus, the catalyst may be overheated, which causes its sintering and an irreversible loss of its activity. At the same time, some parts of the catalyst bed may be reactivated to a lesser extent, e.g. as the result of nonuniform flow of the reactivating medium. The above disadvantages of the reactivation of catalysts in deep layer in the reactor have initiated the study of reactivation outside the catalytic reactor (offsite reactivation). In this process, the catalyst can be placed in a shallow layer (3-5 cm high) on a perforated grate through which reactivation air is introduced with programme-controlled temperature. In the shallow catalyst layer the danger of catalyst overheating can be excluded and the process of coke-combustion can be reliably controlled. On discharging the catalyst from the reactor, mechanical impurities and crushed pellets can be removed which otherwise increase the hydrodynamic resistance of the bed. The aim of this work was to determine the shallowbed reactivation conditions (flow rate of reactivation air and the rate of its temperature increase) which would ensure the reliable control of catalyst temperaFormation

ture, thus avoiding catalyst overheating. For this purpose, the model of shallow layer catalyst reactivation had been proposed which made it possible to calculate temperature in the catalyst layer during reactivation. The model is based on the assumption that coke-combustion takes place in the whole catalyst particle uniformly and includes the effect of external heat transfer. The model can be used to predict the conditions under which a fast rise in catalyst temperature is taking place (catalyst ignition). The situation after ignition cannot be simulated by the model. In this case the process is controlled by oxygen diffusion into the catalyst particle and the model derived for a homogeneous coke combustion cannot account for the jump from the kinetic to the diffusion region.

1717

THEORETICAL

PART

The scheme of the reactivation reactor with a shallow catalyst layer is shown in Fig. 1. The reactivation air is introduced below the catalyst layer and has the temperature Ti, that can be programmecontrolled. If the catalyst bed is thin and broad enough, one can assume that it behaves as an adiabatic ideally mixed reactor. The gas temperature between particles, Tg, is thus the same throughout the bed (see Fig. 1). If the intraparticle heat and intraparticle and external mass transfer resistances are negligible the only transfer resistance in coke combustion is that of external heat transfer between the isothermal pellet and the gas. In the kinetic region the rate of coke oxidation (rc, rs) is first order [l-3] with respect to the amount of carbon and sulphur on the catalyst and to the concentration of oxygen in air. For large excess of oxygen, one can assume that combustion takes place at constant oxygen partial pressure and the rates of carbon and sulphur consumption are given by

rc = kg exp [ - E,/(RTs)]c~Yc

(I)

rs = kg exp [ - Es/(RTs)]c~Ys

(2)

where Y,, Y, are the unreacted fraction of carbon and

1718

KAREL TO

KLUSACEK

et 01.

The shallow bed coke combustion depends on three parameters: the air space velocity, s, the rate of air temperature rise, 4, and the heat transfer coefficient, h. The heat transfer coefficient can be calculated from dimensionless correlations: Nu = /(Re, Pr). For example for particles in packed beds, Whitaker [4] recommends the relation

VENT t

Nu = (0.5Re’j2 where Nusselt, defined as

Reynolds

+ 0.2Re2’3)Pr1’3

and

Prandtl

(10)

number

are

Nu = (Zf.rh&)/(R(1 -&)) Re = (ZapsL)/(

p( 1 - E))

(11)

Pr = cp/A

Fig. I. Batch reactivation reactor.

p and R are dynamic viscosity and thermal conductivity

sulphur in the coke, respectively, c& c8 is the initial mass fraction of carbon and sulphur, respectively, and T, is the temperature of catalyst particle. If we limit ourselves only to linear temperature increase of the air introduced beneath the catalyst bed with time, T,,, then following expression is valid T,,=Tp,+&

(3)

where T % is the temperature of air at the beginning of reactivation experiment (t = 0) and c$is the controlled rate of temperature rise of air beneath the layer. Then, the heat balance in the catalyst particles, with the ratio of volume to outer surface VP/S, = a/3, has the form - c,(dT,ldt)

= rcqc + rsqs + 3V,

--Ts)l(p,a) (4)

where cps and p, are the specific heat and apparent density of catalyst, qc, qs are the heats generated by combustion of unit weight of carbon and sulphur deposits, respectively, h is the heat transfer coefficient between the particle and the gas and a is the equivalent radius of a spherical particle. Heat balance for the gas in the bed voids can he formulated as (dT,ldt)

= (s/ML

--T,) + (1 - s)3V’,

--T,)l(sc,~) (5)

where s is the air space velocity (ratio of volumetric air flow rate under ambient conditions to the volume of catalyst bed), E is the bed porosity, p the air density under conditions of flow measurement and cp the specific heat of air. The mass balances of carbon and sulphur in the catalyst pellets are given as

of air, respectively and L is the height of the catalyst bed. The assumption of ideal mixing in the bed ceases to be valid with increasing bed height, L, when the temperature of particles and gas increase in the air flow direction. This can be included into the model by replacing one adiabatic reactor with a cascade of such reactors (the number of reactors can be taken equal to the number of layers of catalyst pellets). Preliminary calculations have shown that numerical integration of the complete dynamic model is quite time-consuming and simplification of the model is desirable. It has appeared that the accumulation terms of eqs (4) and (5) are very small in comparison with other terms of the heat balances. Differential eqs (4) and (5) can be, therefore, replaced by algebraic relations Tg = T, + [(I - s)PJ(cpPs)](rcqc T, = Ts +

+rsqs)

C(w,V(3~)1(r&k + rsqs)-

(12)

(13)

The simplified model is, thus, represented by two differential mass balances, (6) and (7), and two algebraic eqs (12) and (13). Initial conditions for differential eqs (6) and (7) have the form (9). Comparison of results shows only insignificant differences in T,(t), T,(t), Ye(f) and Ys(t) between the complete and simphfied models. Provided that the temperature rise of introduced air, T,,, is linear with time [eq. (3)], the calculated time profiles of carbon and sulphur [ Y,(t), s(t)] as well as of temperatures [T,(t), T,(t)] can be easily transformed into dependence on the inlet air temperature, Ti,.

EXPERIMENTAL

PART

Catalyst

c‘$(d Yc/dt) = - rc,

(6)

c,o(d Y,/dc) = - rs.

(7)

The dynamic model of. catalyst reactivation [eqs (4k(7)] is complemented by initial conditions

t=O;

T,=T,=T;

(8)

t=o;

Yc=Ys=l.

(9)

In reactivation experiments, an industrial deactivated cobalt-molybdenum catalyst was employed that has been used in hydrodesulphurization of primary oil products (light petroleum to diesel oil). The catalyst pellets were cylindrical (diameter x height = 7.5 x 6.5 mm). After 15 months of continuous operation, the catalyst contained 10 wt. ok carbon and 3.7 wt. y0 sulphur and its activity was 40% of the initial value.

HDS catalyst regeneration

plemented by initial conditions (9) and two algebraic equations (12) and (13). Integration of the set (6) and (7) was performed by the Runge-Kutta method. Temperatures Tg and T, were calculated during integration from eqs (12) and (13), using an iteration method. The parameters used in calculations are listed in Table 1. Kinetic parameters of reactivation (Ei, ko; i = C, S) were determined experimentally recently [5]. Figures 2 and 3 illustrate the results for two typical cases. In Figs 2(a) and 3(a) the linear rise of inlet air temperature, Ti,, is shown together with the calculated temperature rise in the bed, T,. These results are presented in coordinates ATr vs Ti, in Figs 2(b) and 3(b), together with experimental results. It can be seen that the model describes the actual temperature conditions during reactivation well. Figure 2 documents the gradual coke combustion of deposits. The situation in Fig. 3 corresponds to conditions under which catalyst ignition takes place, i.e. a fast temperature rise by several hundred degrees. Reactivation conditions for both cases differ only in the rate of air temperature rise, 4, the air space velocity, s, is identical for both cases. It is seen that for high rate of air temperature rise, & = 80 K h- ’ (Fig. 3), the heat generated by exothermic coke oxidation is not completely removed by the air. The air temperature in the bed and the pellet temperature increase and the rate of combustion accelerates. This results in ignition of the catalyst and in switching to diffusion limited kinetics: the simulation cannot go, therefore, any further. The ignition temperature was around 350°C for all combinations of parameters sand 4 used. Around 360°C the experimental dependence

Reactivation reactor Reactivation experiments were performed under atmospheric pressure in the reactor shown schematically in Fig. 1. The reactor construction made possible programmed increase of temperature of the air introduced beneath the catalyst bed. The reactor was made from stainless steel tube (1) (id. 6.8 cm, height 30 cm). The upper part of the reactor in which catalyst pellets were placed was 12 cm long. To ensure adiabatic conditions the reactor was isolated by ceramic wool (2) (10 cm). Reactivation air was introduced via a flow meter (3) into the bottom part of the reactor in which it was heated by resistance heating (4). The heating was controlled by controller (5) connected to a temperature programmer (6). The heated air passed first through a 3 cm high equalization bed of ceramic beads (7) placed on a perforated grate and then through the catalyst bed (8). The height of the catalyst bed ranged from 1.5 to 4 cm. The temperature of air just below the catalyst bed (Tin) was followed by thermocouple Tl. Preliminary experiments showed that radial air temperature differences below the bed did not exceed 10°C. The temperature of combustion products leaving the bed (T,) was read on thermocouple T2. Reactivation experiments For shallow catalyst beds, the temperature conditions in the bed are determined by two parameters: by the space velocity of air, s, which determines the heat transfer coefficient h [see eqs (10) and ( 1 1)] and by the rate of air temperature rise, 4. The aim of the experiments was to determine the effect of these parameters on reactivation and to find the conditions for slow coke combustion without danger of catalyst ignition. Experimental results were compared with the calculated results [eqs (6) (7), (12) (13)]. For evaluation of experiments, the most illustrative is the dependence of air temperature rise across the bed, AT, = Tg - T,,, on the inlet air temperature, T,,. The difference ATs was determined by thermocouples T2 and Tl (see Fig. 1). RESULTS

1719

Table 1. Parameters for numerical simulation (simplified model) I

= 5 x 10m4 Js-

p

= 3.28 x 10m4 gem-’

‘cm

‘Km’

I?,=

s-l

Es = 48.37 kJ mole’ kg = 1.23 x 105s-’

= 1.07 Jg-‘K-l

CP p =

1.2 x 10e3gcme3

a = 0.357cm

AND DISCUSSION

c

k” s = 25.6 s- ’

qc = 32.8 kJ g- ’ =41.9kJg-’

p, = l.25‘gcme3

The model of shallow bed catalyst reactivation contains two differential equations (6) and (7) com-

T3

= 0.42

0

0

4

m T,n.'=C

t,h

Fig. 2. Reactivation CES 40:sT.

without

ignition

(s = 30,000

h-‘,

110.37kJmol-’

4=4OKh-‘)AT,vst(a);AT,vsT,,(b).

1720

KAREL

2

KLUSACSEKet al.

4

100

3co

T,”.OC

t.h

Fig. 3. Reactivation with ignition (s = 30,000

falls down rapidly [Fig. 3(b)]+atalyst extination. Approximately at 500°C the heat evolution is stopped (AT’ = 0), i.e. the coke is completely burned-off. At lower rate of air temperature rise, 4 = 40 K h- I, (Fig. 2) the catalyst ignition does not take place. The curves ATE---T& show two distinctly separated peaks. The maxtmum of the first (smaller) peak is around 225°C ( ATs = 12”C), the maximum of the second one is about 38%39O”C (AT8 = 35°C). In all experiments without ignition the AT, vs Tin dependences were similar. It is, therefore, apparent that during coke oxidation, two partly overlapping consecutive processes take place: the oxidation of sulphur and carbon. Figure 4 shows the calculated fractions of unreacted carbon ( Yc) and sulphur (Y,) on the catalyst. The air space velocity, s, and the rate of inlet air temperature rise, 4, are the same as in Fig. 2. Figure 2 demonstrates that at about 150°C combustion of sulphur is fast (Ys rapidly decreases). Sulphur combustion is essentially complete at 30%31O”C, after that carbon is combusted as evidenced by the steep decrease of Y,. The rate of heat generation is substantially greater for carbon than for sulphur combustion. That is why ignitions were usually caused by the second exothermic process (see Fig. 3). Only for very fast temperature rises, 4. and for low air space velocities, s, ignition was observed also during suIphur combustion. The main requirement for successful catalyst reactivation is the safe process control with catalyst ignition being avoided. To find the conditions for gradual coke combustion, we performed numerical simulation of the reactivation of the spent HDS catalyst for different combinations of air space velocities, s (corresponding to Reynolds numbers in the range 10-350) and inlet air temperature rises, &. The results are shown in Fig. 5, together with experimental points obtained in the laboratory batch reactivation reactor. On the basis of simulation results the plane of parameters s and +J can be divided into two regions separated by the broken line in Fig. 5. The combinations of parameters s and 4 that lie above this line have led always to catalyst ignition during simulation, while the region below the line ensured gradual combustion without ignition.

I 500

.O

h- ‘, C#J = 80 K h - ’ ) ATg vs t (a); ATE vs Tin (b).

Fig. 4. Unreacted fractions of sulphur ( Ys) and carbon (Y,).

0

10

30

s

lo-‘.

50 h-’

Fig. 5. Parameters for safe reactivation (a) and ignition (b). Experimental points: 0 with ignition, 0 without ignition.

Figure 5 also demonstrates good agreement between experimental and simulation results for points that lie very close to the boundary. To perform reactivation in the vicinity of the boundary cannot be, however, recommended because the jump into the ignition region can be induced by small variations of conditions, e.g. small changes in air space velocity. Safe conditions for HDS catalyst reactivation are, e.g. s = 3 x lo4 h-’ and 4 = 40 K h-’ for which the results in Fig. 2 are valid.

HDS catalyst regeneration CONCLUSIONS

The model of shallow bed reactivation of a cobalt-molybdenum HDS catalyst with excess air has been proposed. Comparison of calculated results with experimental data obtained in a batch reactivation reactor shows that the model is able to predict reliably the conditions for catalyst ignition. Reactivation conditions (s, 4) under which reactivation is slow and can be safely controlled, were determined. The model can also be utilized for other catalysts and adsorbents. The model can also be easily extended to a nonlinear rise of inlet air temperature and to a continual reactivation process where the catalyst is placed, e.g. on a stainless-steel band which passes through a tunnel furnace along which the suitable temperature profile is maintained.

NOTATION

a CO

Cp’ cp.5

E, k* h L 4 r

3Vp/Sp; equivalent pellet radius (ratio of volume to external surface) initial mass fraction specific heat of air and catalyst, respectively =

constants of Arrhenius equation heat transfer coefficient between gas and pellet height of catalyst bed heat generated by combustion of unit weight of carbon or sulphur deposits rate of combustion

1721

gas constant space velocity of air outer pellet surface time inlet air temperature gas temperature in the bed catalyst particle temperature

R S

=Tg-q, = Ts-Tg

pellet volume fraction of unreacted carbon and sulphur, respectively bed porosity thermal conductivity dynamic viscosity air density apparent pellet density rate of inlet air temperature increase Subscripts

carbon sulphur

C

S

REFERENCES

Satterfield Ch. N., Mass Transfer in Heterogeneous Catalysis. MIT Press, Cambridge, Mass. 1970. [Z] Hano T., Nakashio F. and Kusunoki T., J. them. Engng

[l]

Japan [3] [S] [5]

1975 8 127-130.

Weisz P. B. and Goodwin

R. D., .I. Catal. 1966 6 227-236. Whitaker S., A.I.Ch.E. J. 1972 18 361-371. Fott P. and &estaP., CON. Czech. Chem. Commun. 1983 48 334&3355.