from Waste Gases by Reaction with MnO

43 Removal of SO from Waste Gases by 2

Reaction with MnO on Gamma-Alumina x

P.

J.

W.

M.

VAN

DEN

BOSCH

and

W.

A.

DE

JONG

Downloaded by CORNELL UNIV on May 18, 2017 | http://pubs.acs.org Publication Date: June 1, 1975 | doi: 10.1021/ba-1974-0133.ch043

Laboratory of C h e m i c a l Technology, Delft University of Technology, Julianalaan 136, Delft, T h e Netherlands

Sulfur dioxide can be removed from oxygen-containing waste gases by reaction with MnO on γ-alumina between 300° and 600°C. Since reduction of the resulting Mn-sulfate with hydro gen is reasonably rapid at 475°C, isothermal swing operation with a fixed-bed reactor system is theoretically possible. This paper deals with the complex"pseudo-autocatalytic"reduction mechanism. Experiments in a thermobalance gave kinetic data which were used in calculations with a mixed homogeneous/ heterogeneousfixed-bedreactor model. The results of the calcu lations agree semiquantitatively with experimental data obtained on a fixed bed containing Mn-sulfate on γ-alumina. x

aste gas desulfurization is one way to eliminate sulfur dioxide pollution i n flue gases produced by burning sulfur-containing fuels and i n stack gases of smelters and sulfuric acid plants. Several methods for sulfur removal from stack gases are based on simul taneous reaction of S 0 and 0 w i t h a solid sorbent consisting of or containing an oxide from w h i c h a stable sulfate can be produced. Since such processes should be regenerative to avoid secondary soil or water pollution, the sulfate must be decomposed thermally or chemically. If reductive regeneration can be done at about the same reactor inlet temperature as the sulfation, fixed-bed reactors can be used, thus avoiding attrition problems associated w i t h pneumatic transport of the solid ( I ) . Such a fixed-bed process is the Shell F G D , i n w h i c h the reactive solid is copper oxide on alumina ( 2 ) . W e expected that manganese oxide w o u l d also be a suitable agent for S 0 removal because the literature indicates that simultaneous sorption of S 0 and 0 by M n oxides is possible. Calculations suggest that the reduction of M n sulfate w i t h H is thermodynamically possible at moderate temperatures. Since preliminary experiments indicated that sulfation and reduction w i t h H both proceed at acceptable rates w i t h MnO^, on γ - Α 1 0 at 475°C, a program was set up to establish whether the following cyclic regenerative process is feasible: ( 1 ) Sorptive reaction of S 0 and 0 w i t h MnO^. on a highly porous carrier w h i c h provides a large surface area as w e l l as m i n i m u m diffusional resistance to the reaction;

W

2

2

2

2

2

2

2

2

2

3

2

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Hulburt; Chemical Reaction Engineering—II Advances in Chemistry; American Chemical Society: Washington, DC, 1974.

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CHEMICAL

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(2) Reduction of the sulfate formed on the carrier w i t h H , giving regen erated solid and a gaseous product w i t h a h i g h sulfur content. 2

Literature Data on Process Chemistry General. M a n y sorption/regeneration cycles can be done w i t h MnO / A 1 0 . T h e active oxide from w h i c h the sulfate is formed is a complex mixture of M n oxides containing M n O , formed by oxidation of initially formed M n O . A summary of a literature survey of the chemistry of sulfation and regeneration is given below. Sorption of S 0 / 0 . Three groups of papers on the sulfation of M n O and M n 0 by S 0 and 0 can be distinguished. I n the first are the purely chemical studies—e.g., H a m m i c k (3) and Davis (4) reported complex solid intermediate products, irrespective of the starting material, and agreed that the final product was invariably M n S 0 up to 800°C. Above 800°C the sulfate decomposes thermally. T h e second group contains kinetic studies of sulfation and oxidation of M n oxides. Ingraham and Marier (5) studied sulfation i n a thermobalance at partial pressures of S 0 and 0 m u c h higher than those i n stack gases—viz., 0.33 and 0.66 atm. Unfortunately, their data cannot be used here. I n the third group are studies (6, 7) on the use of M n oxides i n stack gas desulfurization. One of these describes (6) the use of fixed-beds of M n oxides for S 0 removal between 130° and 330°C. A coprecipitate of M n and A l hydroxides is recommended by A V C O C o r p . (8) as a sorbent i n a regenerative system. However, their experimental data do not agree w i t h thermodynamic calculations; these predict that M n S should be produced upon reduction of M n S 0 by H , but M n O and M n S are actually found, the M n O content being unexpectedly high. U n o et al. (9) describe the D A P - M n process, i n w h i c h S 0 is sorbed by an uncommon M n oxide of general formula M n 0 · n H 0 · ( O < i < 1 and Ο < η < 1 ) . This solid, w h i c h is formed i n a regeneration step consisting of precipitation of M n hydrox ides w i t h ammonia from M n S 0 solution, filtration, and subsequent oxidation, is very reactive towards S 0 as compared w i t h various other M n oxides. Graham (10) studied the performance of a fixed bed reactor for the reaction of S 0 w i t h M n 0 to give M n S 0 . A recent patent to Kennecott Copper C o r p . (11) describes the desulfurization of oxygen-containing stack gases w i t h solids containing M n O . Results show h i g h conversion of the solid ( 1 0 - 3 0 % ) between 100° and 400°C. Reduction of Sulfate. Reduction of M n S 0 w i t h various agents has been studied by Fuller and E d l u n d ( 1 2 ) , who found conversion to M n O and M n S w h i c h depend on temperature and amount of reducing agent. T h e i r proposed mechanism involves reaction to M n S 0 and two possible subsequent reactions —viz., decomposition to S 0 and M n O and disproportionation to M n S and M n S 0 . T h e theory fits w e l l the data of C o l a (13) who studied the reaction of M n S 0 upon heating i n a stream of N and i n an autoclave. I n the autoclave M n S and M n S 0 were formed, no solid desulfurization taking place, but i n the flow experiment an increasing amount of oxide was formed w h e n the tem perature was higher. T h e above literature indicates a temperature of at least 600°C for reasonably fast and complete sulfate reduction. A l t h o u g h this is high, regeneration is feasible because the sulfate is stable i n an oxidizing atmosphere u p to 800°C. Oxidation of Acceptor. W h e n oxygen-containing gases are passed over a regenerated ΜηΟ,,./γ-Α1 0 acceptor, the solid is oxidized and a large amount w

2

3

2

2

2

3

a

2

2

2


4

2

2

2

4

2

2

1+ i

2

4

2

2

2

4

4

3

2

4

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of heat is generated. Consequently, deactivation and sintering can occur. A m o n g the possible oxidation reactions are those of M n S and M n O produced by previous sulfate reduction. Bourgeois et al. (14) quantified such heat effects in a modeling study of a fixed-bed reactor containing C u 0 on alumina. T h e y found that a temperature peak as h i g h as 100°C above the inlet temperature travels through the bed during oxidation to C u O . T h e oxidation of M n O has been studied by Moore et al. (15) and Pechkovskii et al. (16). T h e first group found M n 0 as the solid product whereas Pechkovskii detected also M n 0 . Although M n 0 is thermodynamically the most stable oxide at the conditions applied, the formation of this compound is not reported by anyone. Roasting of M n S was studied by Batsonov et al. (17) and Chagunava et al. (18). Appreciable desulfurization of the solid is observed if the temperature is kept below 980°C. Conclusions from the Literature Study. Most studies deal w i t h unsup ported M n compounds (except Ref. 8) ; we used a disperse system on γ-alumina, w h i c h may influence the course of some reactions. F r o m the literature it appears that M n S 0 is likely to be the main, if not sole, product of the reaction between MnO^-containing acceptors, S 0 , and 0 . T h e reactions occurring during reduction w i t h H are probably very complex (see Table I ) . 2

2

3

3

4


2

4

2

2

2

Table I. I II III IV V VI VII VIII IX X X I a. b. XII

Mn 0 2Mn 0 6MnO 4Mn 0 MnS 2MnS MnS0 MnS0 3MnS0 MnS 3MnS0 3MnS0 S0 3

4

2

3

3

4

+ 3 S 0 + 0 -> 3 M n S 0 + 4 S 0 + 0 -> 4 M n S 0 + 0 -> 2 M n 0 + 0 -» 6 M n 0 + 20 — MnS0 + 30 -> 2MnO + 2 S 0 +4H -»MnS + 4H 0 + H -> M n O + S 0 + H 0 + MnS τ± 4MnO + 4 S 0 +H 0 3MnO + 4 S 0 + H 0 + 4H S -> 3MnS + 4 S 0 + 4 H 0 + 3H -> H S + 2 H 0 2

2

4

2

2

4

2

3

2

2

2

4

3

4

2

4

2

2

4

2

Possible Reactions

2

2

4

2

2

4 4

2

2

2

2

2

2

2

2

2

2

2

2

Experimental Materials. Acceptors containing M n compounds were prepared from: (1) γ-alumina 0 0 0 - 1 . 5 E 6 (crushed and sieved to 0.6-1.0 m m ) obtained from Ketjen ( A K Z O ) Amsterdam, surface area 241 m gram" , pore volume 0.7 cc g r a m , silica gel A T 23 (Gembo C h e m i e ) , surface area 470 m gram" , pore volume 0.7 cc gram" , and (2) high purity M n S 0 · 1 H 0 (Merck, D a r m stadt). The acceptors were prepared by impregnation methods; i n most cases the procedure i n Ref. 19 was used after modification for M n compounds. Equipment. A C a h n R G - H V thermobalance combined w i t h a PerkinE l m e r microfurnace/programing unit, a modified high-temperature GuinierLamé x-ray camera, and an experimental unit containing a fixed-bed reactor were used. T h e x-ray camera showed changes i n the solid during isothermal reaction. Figure 1 is a flow sheet of the equipment containing two fixed-bed reactors, quartz tubes of 8 m m i d and 300 m m long. Reactant gases were metered by a combination of Brooks E L F flow controllers ( 1 ) and Fisher-Porter Rotameters ( 2 ) . Electromagnetic valves (3) were operated by a timing mechanism to control the complete cycle and feed gas composition for reactors R l and R 2 . Water was fed by special motor buret (4) i n combination w i t h electromagnetic valves. T h e gases were preheated, and water was evaporated i n preheaters 2

1

-1

2

1

4

2


1


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CHEMICAL

Figure 1.

REACTION

ENGINEERING

II

Twin-reactor flow equipment

E l and E 2 . Samples of product gases were injected automatically into an on-line gas chromatograph ( 5 ) ; a fully automatic acidimétrie titration (6) i n 5% H 0 solution was also used to monitor product gases. I n the G S C system the components H , H S , S 0 , and H 0 were separated at 130°C over a 3-m Chromosorb 104 column (7) using H e as the carrier gas. 2

2

2

2

2

ι

Figure 2.

time

2

(min)

Influence of temperature on Reaction VII (Phydrogen =

0.73

Otm)

Experimental Process C y c l e . The reactor was filled w i t h about 5 grams of acceptor. If the acceptor was i n the sulfate form, it was heated to 475°C while passing through N ; when this temperature h a d been reached, regenera tion was started b y introducing H . After 50 m i n regeneration, the reactor was flushed w i t h N ; then a simulated flue gas was introduced w h i c h contained 2

2

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known amounts of S 0 (say 0 . 4 % ) and O . Sulfation was continued for 6 0 m i n , the reactor was again flushed w i t h N , and a second cycle was started. Since almost all the sulfur is usually converted to S 0 upon regeneration, continuous monitoring of the gases leaving the reactor indicated whether sulfur recovery was complete. The high degree of automation allows life testing of acceptors; more than 1 0 0 0 cycles were r u n w i t h a promising acceptor. 2

s

2


2

0

5

Figure

3.

15

10^

time

20

25

(min)

Effect of temperature on (pH s = 0.05 atm)

Reaction

XI

t

Results on Process

Chemistry

Sorption of S 0 . Thermobalance experiments (Figures 2 and 3 ) , x-ray data, and flow experiments prove that i n the sorption phase of the cycle, sulfate is formed on the acceptor at 4 7 5 ° C according to: 2

2 M n 0 + 4S0 + 0 2

3

2

2

-> 4 M n S 0

4

O n a basis of mass balances we established that M n 0 is the active oxide, but we are not sure about its exact crystalline form. Examples of S 0 breakthrough curves obtained i n the flow equipment are given i n Figure 4. The amount of M n on the acceptor seems to have a marked influence on oxide conversion under dynamic conditions; presumably this results from diffusional retardation i n the solid product layer, its density being smaller than that of the solid reactant. Similarly incomplete conversion was obtained when sulfating pure M n 0 and MnO^. on γ-alumina i n the thermo balance. Calculations show that rate of sulfation of MnO^. on γ-alumina is also limited by pore diffusion, w h i c h necessitates an acceptor support of h i g h porosity and surface area on w h i c h the M n O ^ is very finely dispersed. It was also found that the temperature, space velocity, and water content of the simulated flue gas have very little influence on acceptor capacity. Reductive Regeneration. The thermobalance experiments of Figure 2 indicate that the temperature should exceed 5 5 0 ° C when regenerating w i t h H . T h e m a i n product is expected to be M n S , indicating that desulfurization of the solid is incomplete. Thus, it is surprising that H reduces M n S 0 on 2

3

2

2

3

2

2


4


Figure 4.

SO breakthrough with various acceptors at 475°C (pso = poj = 0.05 atm) s

g

0.004 atm;

γ - Α 1 0 rapidly even at 475°C, w i t h fairly complete conversion of the solid sulfur compound into (mainly) S 0 (Figures 5, 6, and 7 ) . However, regenera tion at 475°C is fast only when γ-alumina is used as the support for M n O ^ : neither unsupported M n S 0 nor M n S 0 on silica can be reduced w i t h H at acceptable rates at 475°C; at least 600°C is required for this reaction. These facts as w e l l as the negative effects of water on the rate of regeneration (Figure 6) are explained as follows: presumably the active species during regeneration w i t h H is H S , formed initially b y hydrolysis of traces of M n S originating from the slow reduction of M n S 0 w i t h H : 2

3

2

4

2

4

2

2

4

2

M n S 0 + 4 H -> M n S + 4 H 0 4

2

2

MnS + H 0 -* MnO + H S 2

2

80-

- 60 60 -

H O 2

0

20 m

time

40

60

80

(min)

Figure 5. Conversion to S 0 and product composition during acceptor regeneration (Ύ = 475° C, GHSVE = 120 hr' ; previous sulfation with dry gas) 2

2

1


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Reduction of M n S 0 b y H S is m u c h faster than w i t h H . T h e product is S 0 : 4

2

2

3 M n S 0 -f H S 4

2

2

3MnO + 4 S 0 -1- H 0 2

2

I n the fixed-bed reactor, S 0 is formed; further amounts of H S needed for rapid conversion are produced b y hydrogénation of S 0 w i t h γ-alumina as the catalyst. Figure 8 shows that hydrogénation of S 0 is possible over γ - Α 1 0 ; 2

2

2


2

0

20 time

40

60

2

80

(min)

Figure 6. Conversion to SO and product composition during acceptor regeneration (conditions similar to those of Figure 15 but previous sulfation with wet gas) s

100

Figure 7. Conversion to SO ζ and product composition during acceptor regeneration (T == 525°C; GHSV = 120 hr' ; previous sulfation with dry gas) Ht

1


3


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Figure 8. Hydrogénation of SOg over y-alumina (450°C, GHSV = 6000, p * = 0.003 atm) so

silica activity is quite small or negligible under conditions where γ-alumina is an effective catalyst. W a t e r supposedly retards the hydrogénation of S 0 , a phenomenon also found w i t h the Claus reaction (27). Thus, the reactions w h i c h occur during regeneration w i t h H show ' pseudo-autocatalytic" behavior. T h e initial slow attack b y H on M n S 0 is followed by a rapid conversion of the sulfate w i t h H S , formed in situ from H and S 0 , the gaseous product of the reaction between M n S 0 and H S . T o strengthen this qualitative and intuitive picture of the regeneration, we set up a model of a non-stationary fixed-bed reactor to simulate the reduction of M n S 0 on γ-alumina. 2

2

2

4

2

2

4

2

2

4

Modelling

of Regeneration

in a Fixed-Bed

Reactor

M o d e l Equations and Assumptions. A mixed homogeneous/heterogeneous reactor model was derived. W e feel that Reaction X I a predominates over Reaction X l b w h e n regenerating i n a fixed bed. The two sets of rate equations included i n the model are listed i n Table II. T h e mathematical model contains six homogeneous mass balances: three for gases, three for solid compounds, and one heterogeneous mass balance for water. T h e contribution of the solid phase—i.e., of adsorbed gases, to the mass balance for H , H S , and S 0 is negligible, but this is not true for water, w h i c h is adsorbed i n appreciable amounts. P l u g flow and isothermal operation are assumed, and transport re sistances are considered absent. T h e mass balance for H , H S , and S 0 is as follows: 2

2

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2

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T h e mass balance for the solid components, M n S 0 , M n O , a n d M n S , i s : 4

17 dt

(2)

= Σ vu n ,

For water we have two mass balances: dpmo

U (-IT)

+

dl

V~j

-

-

( '°" C

( Σ VU H 0 . i r 2

g

t

f

g (

0.ij +

H 2

0 and

t

= 0:

PH

0 and

t

H2

2

= 0 and * > 0:

=

= 0:

*H 0 2

(3)

2

= 0, PH O =

2

j

" " ) (PH O) (1 - 9H O)

= #ad.PH 0 (1 - 9 o) for

[

(Xde.Sn o)

0, p

(4)

2

H 2

s = 0, pso = 0 2

1, PH O = 0, P H S 2

2

= 0, pso - 0 2

= 0» PH O = 0, P H S

= 0, pso - o,

= 1, C n n O

~ 0, CMnS - 0

2

2

ΟΜΠΘ0

4

2

These equations for the reduction are formulated as if the systems were homogeneous. This is permitted only w h e n the reactive ingredient is so highly dispersed o n the carrier that mass transfer limitations through the product layer can be neglected, w h e n diffusion i n the pores of the carrier does not limit the rate, and when the particles of the reactive ingredient of the acceptor react homogeneously. These conditions appear to be fulfilled for the reduction of M n S 0 o n γ - Α 1 0 . T h e above partial differential equations were solved b y the method of characteristics described i n Ref. 20. 4

2

Table I I .

3

Reactions and Rate Equations U s e d i n Fixed-bed Simulation

Set 1

VII XII XIa X

M n S 0 + 4 H - » MnS + 4 H 0 S 0 + 3 H -> H S + 2 H 0 3 M n S 0 + H S -> 3MnO + 4 S 0 + H 0 MnO + H S -* MnS + H 0 4

2

2

2

4

r r = r=

2

2

2

2

2

2

2

H2

2

H2

3

M

4

2

H2

n

*H,O)

4

2

Set 2 same as set 1 but Reaction X I a replaced by : Xlb 3 M n S 0 + 4 H S τ± 3MnS + 4 S 0 " + 4H 0 4

fc,p C nso fc p Pso (1 fc 3o sCM so

2

2

r=

fc p sC nso ?

H2

M

4

2

K i n e t i c D a t a from Thermobalance Experiments. G E N E R A L A S P E C T S . C a l culations w i t h the above model require knowledge of the kinetic parameters of the relevant reduction reactions. It is impossible to obtain such data o n the various M n compounds o n γ - Α 1 0 from kinetic experiments i n the thermo balance. Strong adsorption of reactants and products obscures the weight dif ferences from chemical reactions so m u c h that reliable conversion data cannot be determined. Another problem is the adsorption of water o n the carrier w h i c h occurs under reaction conditions. This is more serious because water retards the hydrogénation of S 0 to H S over γ - Α 1 0 , the key step w h i c h governs the overall rate of acceptor regeneration. Since the flow regime i n the thermo balance differs markedly from that i n the fixed b e d and since this regime is one of the factors w h i c h determine the water concentration o n the surface, the rate of acceptor reduction i n the thermobalance also differs from the rate i n the fixed bed. 2

2

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3


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T h e only way to overcome these difficulties is to determine the kinetics of the reduction reactions involved from thermobalance experiments on the unsup ported M n compounds. This poses the problem of converting the data on such model compounds into rate expressions v a l i d for the disperse systems actually used i n the process cycle because the dimensions of the reactive particles are vastly different i n the two cases. T h e extent to w h i c h mass transfer limitations occur i n thermobalance and fixed bed are expected to differ. This problem can be tackled as follows. T h e rate of the elementary reac tion between solid and gas is taken to be first-order i n each of the two reactants. If the concentration of the gas is expressed by partial pressure ρ and that of the solid i n m o l e s / k g solid reactant, the rate equation becomes:


r

w

= k ρ C w

(5)

B

In the "shrinking core m o d e l " for gas/solid reactions the dependence of the rate on the concentration of the solid reactant is neglected; nevertheless, this model holds quite w e l l for a number of gas/solid systems because the reaction proceeds i n a relatively thin zone, whose thickness is small compared w i t h the particle diameter. This zone does not change i n thickness over a relatively wide range of solids conversion as it travels inwards. I n terms of rate E q u a t i o n 5 this implies that C is constant, and the rate equation simplifies to: s

r in which C

s

0

w

= k pC , ο w

(5a)

e

is equivalent to the reciprocal of the molecular weight: Τ w.pure solide = k ν,φΜ'

1

(5b)

However, w i t h very small particles such as those i n disperse systems like our acceptors, C can no longer be considered constant. T h e reaction zone then is of the same order of magnitude as the dimensions of the reactive particles, w h i c h , i n other words, react more or less homogeneously (21). Therefore, Equation 5 should be used i n the regeneration model. Nevertheless, the basic kinetic parameter, fc , can be found from experi ments on solids w h i c h react according to the shrinking core model. Rate data for this model are usually expressed i n terms of rates per unit surface area r : s

w

8

r = hp

(6)

B

Rate r is related to the rate per unit weight, r , according to s

w

r

w

- rJS

T

(7)

S being the surface area per unit weight i n the reaction zone. Since this quan tity is considered to be independent of the conversion over a wide range (see above), it should also apply to the initial reaction zone—i.e., the outer surface S , and it can, therefore, be calculated from the diameters of the particles used i n the kinetic experiments. T h e above approach is similar to that followed b y Schwab (22) i n a study of gas/solid reactions; it was also used by Mars (23) i n his study of the C O shift reaction on iron oxide catalysts. Mars related the frequency factor to the number of collisions of gas molecules w i t h the surface, assuming that each collision leads to reaction. So far, diffusion effects have not been discussed, although i n many cases diffusion resistances i n the solid reactant and the product layer can affect the r

0


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overall rate of solid conversion. Models applicable to such cases have been developed by Shen and Smith (24) and by Szekely et ai. (25, 26). These authors also give criteria for establishing the possible occurrence of diffusion limitations. MODEL APPLIED TO ANALYZE T H E THERMOBALANCE DATA. T h e criterion developed by Szekely (25, 26) was applied to show that intraparticle diffusion resistance affects the rates of the gas/solid reactions under the conditions pre vailing i n the thermobalance. A similar conclusion was reached from calcula tions based on the criteria of Shen and Smith (24): it appears that diffusion resistance cannot be neglected. Thus, we used one of the models of Shen and Smith (24)—i.e., the model incorporating diffusion i n the particle and chemical reaction but no film diffusion, to describe the experimental results obtained i n the thermobalance. T h e relevant equation is: θ =

b f c , C

'' ^ p

M a 8 0 <

from Waste Gases by Reaction with MnO

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