tests. The findings of the on-line control and optimization studies are presented in Table I1 and Figure 1. Discussion
The R E F L E X method of optimization apart from the overall on-line program performs efficiently and rapidly in the presence of experimental error. Graphical and tabular results of both on-line and off-line optimization studies have been noted. The off-line explicit function optimization tests of the REFLEX performance reflect the ability of the method to deal concisely with optimization problems of a n empirical nature. The use of the more common test functions for comparative studies indicates the reliability of the R E F L E X to perform on drastically different factor spaces. The results of the on-line tests present more explicitly the nature of the empirical optimization problem. The magnitude of the initial performance indices for all on-line tests are comparable since each is a n evaluation of a simple proportional (digital) controller. The deviation of these performance indices from one another is perhaps a result of many additive effects which were not intentionally introduced into the on-line tests.
The deviations do occur, however, even in these carefully observed studies, and must be accounted for by the optimization scheme if optimization is to occur. Thus, the optimization effort of the REFLEX, when begun under identical conditions, improved the control effort (see Table 11) as per the performance index from approsimately -0.61 to approximately -0.01 regardless of the number of controller parameters employed per test. literature Cited
Box, G. E. P., Applied Statistics, 6, 3-22 (1957). Glass, R. W., LIS Thesis, Clemson University, Clemson, S.C., 1967. Glass, R. R., PhD Thesis, Clemson University, Clemson, S.C., 1970. Nelder, J. A,, Mead, R., Computer J.,7, 308-12 (1965). Spendley, Pi.,Hext, G. R., Himsworth, F. R., Technometrics, 4, 441-61 (1962).
RECEIVED for review June 10, 1971 ACCEPTED August 10, 1972 Work supported by a Xational Aeronautics and Space Administration Traineeship and the Chemical Engineering Department of Clemson University.
Removal of Sulfur Dioxide from Stack Gases by Catalytic Reduction to Elemental Sulfur with Carbon Monoxide Robert Querido' and W. Leigh Short2 Department of Chemical Engineering, Cnivvsity of Massachusetts, Amherst, Jiass. 01002
A process i o catalytically reduce sulfur dioxide in stack gases to elemental sulfur is described. Experimental results indicate that a maximum sulfur compound removal efficiency of alumina catalyst.
P o l l u t i o n of the atmosphere by sulfur dioxide is rapidly becoming a serious ecological problem. The fossil-fueled power plant is the chief source of the sulfur dioxide emissions accouiit'ing for nearly 50% of the total (Hangebrauck and Spaite, 1968). The remaining emissions emanate from smaller, widely separated sources including industry, space heating, motor vehicles, and incineration. Although the control of sulfur oxide emissions from the smaller sources should not, be ignored, almost all of the emphasis in SO2 abatement technology has been directed toward t'he power plants which represent large localized sources. This t'echnology falls into four major categories: (1) Change the energy source to hydroelectric or nuclear. ( 2 ) Switch t o a low-sulfur fuel. However, the availability of this fuel is already critically short and expensive so that
Present address, The Firestone Tire & Rubber Co., Central Research Laboratories, Akron, Ohio 44301. To whom correspondence .should be addressed. 10 Ind.
Eng. Chem. Process Des. Develop., Vol. 12, No. 1 , 1973
97%
is possible utilizing a copper-on-
universal applicability of this abatement technique may be limited a t this time. (3) Fuel desulfurization. The efficiency of coal desulfurization is still low and prohibitively espensive. Although fuel oil desulfurization can reduce the sulfur content to acceptable levels, the cost of the desulfurized fuel would be substantially increased. (4) Flue gas desulfurization. Although no one flue gas desulfurization process has been developed to the level of universal acceptance, this technique is currently under intensive research. Over i o such processes in various stages of development are currently under study (U.S. Dept. HEW, 1960). The Public Health Service has proposed a set of air quality criteria for sulfur dioxide (1967) which corresponds t o a maximum allowable SO, stack gas concentration of about 200 ppm. However, typical coal-burning power plants produce stack gases containing approximately 2000 ppm (Beinstock et al., 1965). Therefore, the working objective of this study was the development of a process capable of reducing the SO2 stack gas concentration by a minimum of 90%.
Experimental Theory
The process discussed in this paper is the catalyt'ic reduction of sulfur dioxide t o elemental sulfur with carbon monoxide. This process is considered t o possess the following potential advantages: (1) Simplicity of the process. The system consists of a packed bed, preceded and followed by precipitators t o remove fly ash and sulfur by-product, respectively. The process is "one-step," with no cyclic regeneration or recycle streams required. Therefore, no auxiliary equipment is needed to operate on t h e by-product. (2) Elemental sulfur is t'lie direct by-product. ( 3 ) The process is dry. (4) It has been shown that CO will simultaneously reduce SO, under SOn reduction operating conditions (Ryason, 1967). (5) Sufficient CO may be present in the stack gas to react with SO,under proper boiler operating condit'ions. Previous W o r k on SOs Reduction by CO. T h e direct cat,alytic reduction of SOz by reaction wit'h CO
2C0
+ SO2 =
1 ; 2 s 2
+ 2C02
has been studied as early as 1918. However, previous work was performed at relatively high reactant concentrations aiid lengthy residence times. Ferguson (1918) studied t h e equilibrium between CO, COz, SO2, aiid elemental sulfur for initial SO2 arid CO feed concentrat,ions of 257, each, by volume. Lepsoe's study (1940) of SOnreduction included contact times on the order of 1 min and reactant concentrations of about 63y0 CO and 3570 Sos,with alumina as the catalyst. Saxtaxtinskij (1965) studied the reduction of SO, over reduced alunite (64'% Xl?Oaand 13.8y0 FenOs)and a 2 : 1 ratio of CO to
so,.
The scope of t'he stack gas problem differs dramatically from the above investigations. For example, burning a 3.5% sulfur content coal results in a SO?concentration in t'he stack gas of about 0.2-0.2570 by volume (2000-2500 ppm). The gas volume produced by a 1000-mW boiler is on the order of 1.72.0 million SCFM (standard ft3/min) which moves through the equipment a t velocities of 35-40 milesdhr. This flow rate corresponds to contact times (volume of voids1 volumetric flow rate) 011 the order of only a few tenths of a second (Slack, 1967). Some preliminary work at, loiv reactant coriceiitratioii has been reported by Ryason (1967). He observed a maximum SO?removal efficiericy of 977, a t 1 O O O O F a t a contact time of 0.35 aec and a n initial SO? concentration of 2000 ppni. He also sllo\ved simultaneous and virtually complete removal of S O:
KO
+ CO = l/*N? + con
R\.asoii also reported a deleterious side reactioii between CO and S? to form significant quantities of COS. The performaniie criterion used in his study was a n SOyremoval efficieiicy. However, since COS is a toxic gas which is not easily removed, the performaiice criterion in this study has been changed froma SO?removal efficiency t o a total sulfur compound removal efficiency where total refers t o both unreacted SO2 and any COS t h a t is produced. Reaction System. l ' h e three main reaction equilibria bb are: coilbidered iii this proce7-
2C'O
+ SO?
co + co +
1)
= 1/2s2
eS2
1,202
+ 2c01
=
COS
=
CO,
Table I. Stack Gas Compositions. Case A,
Chemical specie
C02 SOx
%
Case 8,
14
0 0
co so2
3.4 0 2 Excess air 18 For combustion of 3.2 wt % sulfur coal. 0 2
a
%
18 0 1.4 0 5 0 2 1
Two addit,ional reactions, which are thermodynamically dependent, but which might be important from a kinetic viewpoint are:
2COS COS
+ SO2
t 3/1?02
= =
3 / 8 2
+ 2COz
con+ SO,
The presence of water vapor in the effluent may merely reduce the partial pressure of all reactants. However, if the catalyst employed also promotes the water-gas shift reaction:
H20
+ CO = Hz + COS
then the presence of Hz may lead t o a myriad of other possible reactions, including:
+ SO2 HsS + 21320 2HzS 2Hy + S2 2HzS + SO?= 2H20 + Sz COS + Hn = CO + HzS COS + H2O COS + HpS 3H2
=
=
Also, sulfur vapor exists in a complicat,ed equilibrium among S 2 , Sa, Sq, S 5 , Ss, S;,and Sa, whose distribution is a function of temperature aiid total sulfur partial pressure (Berkowitz and Chupka, 1964; Berkowitz and AIarquart, 1963; lIeyer, 1965). Therefore, a n additional set of equilibria should be included : Sz =
S,
x
=
3, 4, 5, 6, 7, 8
However, past investigators of systems including sulfur vapor equilibria, on-ing to the lack of considered only the S?-S& accurate thermodynamic data for the remaining species. Their experimental results differed from thermodynamic predictions by about 470 (Gamsoii and Elkins, 1953; l l u n r o and Madsin, 1967). The concentrations of gases iii the stack effluent vary, for example, with the fuel burned, excess air, and boiler efficiency. The composition of a typical stack gas is listed as case X in Table I. I n his st'udy of catalyzed S O reduction with CO, Baker (1965) has shown bhat oxygen if present, preferentially oxidizes CO t o COS. Ryasoii (1967) has shown that the rates of reduct'ion of S O and SO2 with CO are on the same order of magnitude. Since a stack gas contains significant quantities en, sufficient CO must be present t o react with all the oxygen. T o reduce the 2000 ppm of SO?in the stack gas described in Case -1,6.87, CO must be present t o react with the 3.47, 02, and 4000 ppm t o react with t,he SO,. Since there is virtually no CO in the effluent, 7.2y0 CO must be added externally to the stack gas. However, if the excess air m-as decreased from lSyOto about lY0, the 02 content would fall to 0.57, and sufficient CO would be generated to react with all the 02 and SO? (Case 13, Table I ; Perry et al., 1963). Furnace operation a t 1y0excess air will result in some loss in the overall boiler efficiency. The loss in thermal efficiency, due only t o the incomplete combustion of the fuel, Ind. Eng. Chem. Process Des. Develop., Vol. 12, No. 1 , 1973
11
80
60
40
20
0 08
I2
IO
14
16
18
2.0
CO RATIO
08
10
20
data. T h e following species were considered: CO, CO,, SOz, S O I , C O S , CS,,0 2 , S2,S?,and Ss. S o other sulfur species
700° K
O
0 8
M
I O
,
l
12
14
I
,
16
18
1 20
CO RATIO
Figure 2. Equilibrium percent sulfur compounds remaining as function of CO ratio and temperature Case 8, 1 % excess air
amounts to less than 5%. This figure should be reduced since a t low excess air levels t'liere is less air and consequently less wasted sensible heat escaping out the stack. Thermodynamic Calculations. T h e thermodynamic equilibria for both Cases A a n d €3 (18y0 and 1% excess air, respectively) were calculated using the minimization of free energy technique. T h e procedure followed was essentially t h a t of K h i t e (1958). It utilizes a modified steepest descent search. Instead of requiring specification of the important equilibria, the technique requires only the inclusion of all possible reactant and product species. T h e J-ISXF Tables were used as t8hesource of i'ree energy 12
18
Case 8, 1% excess air
loo
0
16
Figure 3. Effects of CO ratio and temperature on equilibrium sulfur dioxide removal
Case A, 18% excess air
Z
14
CO RATIO
Figure 1 . Equilibrium percent sulfur compounds remaining as function of CO ratio and temperature
!
12
Ind. Eng. Chem. Process Des. Develop., Vol. 1 2, No. 1 , 1973
were included in the tables. The temperature range considered in the thermodynamic calculation was 400-1000"K. This represents the range found in a typical coal-burning power plant. The results for both cases indicated that a t thermodynamic equilibrium, SO3 and CS2 concentrations were nil. Figure 1 is a plot of the percent remaining sulfur compounds (unremoved SO,plus COS expressed as a mole percent of the upstream SO?) expressed as a function of CO ratio and temperature for Case A. The CO ratio is defined as:
Rco
3
('Vco - 2NoJ
where S c o , SSO,, and Xo, are the upstream concentrations of CO, SO?,and Or, respectively. This ratio specifies the amount of CO available t o react with SO2 since the reaction between O2 and CO to form C o t is predominaiit (Baker, 1965). The are multiplied by two upstream concentrations, S o 2and XSO? to satisfy the stoichiometry. When this ratio is unity, there is stoichiometrically sufficient CO t o react with all the SO1 and Or present. Figure 2 is the same plot for Case E.Figures 3 and 4 indicate the effects of CO ratio and temperature on the SO, reduction and COS production for Case B,those for Case .I being essentially identical. The follonkg conclusionb can be drawn from these plots: The percent sulfur compounds remaining a t thermodynamic equilibrium has a minimum a t a CO ratio of unity a t temperatures below 800OK. The magnitude of this minimum decreases with decreasing temperature, approaching a value of zero a t 400" K. Thus, from a thermodynamic standpoint, complete sulfur compound removals are possible a t temperatures below 400OK. However, kinetic considerations will determine the lo\y-temperat'ure operating limit. The relatively large decrease in the magnitude of the minimums below 700°K can be attributed to the shift in the sulfur equilibrium. h t 700°K and above, the thermodynamic calculations indicate that S?is the only form of sulfur present.
-- -
G A S BLENDE!
r-------
-
IO00
E a
z
0 -
c 4 a c
z
w
Figure 5. Experimental flow system
6'X 1 / 8 " b r a p a k Q S mlumn Helium carrier - 3 0 m l l m i n
CO2
08
IO
12
14
16
18
2D
021
I
CO RATIO
Figure 4. Effects of CO ratio and temperatue on equilibrium carbonyl sulfide production Case
B, 1 %
excess air
Likewise, a t 500'K and below, Ss is the only form present; whereas a t 600°K both forms exist (Querido, 1971). Thermodynamics indicate t h a t the reaction between CO and Sz is favored over t h a t of CO and SS.Therefore, at temperatures below 600°K where sulfur vapor contains Sa,the COS production decreases and the percent remaining sulfur compounds is decreased. However, S g was not included in the calculation owing to its absence from the JANAF tables. Thus, the actual locations of the minimums are probably less accurate between 5OO'K and 700°K. The minimums become sharper with decreasing temperature; Le., control of the CO ratio would become more critical. A t thermodynamic equilibrium, after the 02 has been reacted to CO,, the SOz and COS concentrations for both Cases A and I3 are virtually identical. However, Case 1 does indicate about 1% better removal which is attributed to the smaller concentration of COZin the feed. Although the SO2 can be completely reduced, thermodynamics indicate the production of substantial quantities of COS.Therefore, one of the major aims of this study was the determination of a catalyst or reaction conditions ivhich were specific to the SO,reduction and which lvould not promote the formation of COS. Experimental Apparatus
Reactor a n d Flow System. T h e reaction system was studied under controlled conditions in the laboratory. Preliminary results were obtained with a floiv reactor operated isothernially and essentially a t atmospheric pressure. T h e reactor was fabricated from 1-in., schedule 40, t y p e 304, stainless steel pipe, 28 in. in length a n d was capable of accommodating 1-12 in. of catalyst (12 in. corresponds t o about 150 grams of metal-on-alumina catalyst). T h e catalyst bed started about 8 in. from the inlet a n d contained 12 Chromel-Alumel thermocouples spaced 1 in. a p a r t a n d located in the center of each 1-in. increment of bed. d Leeds & S o r t h r u p millivolt potentiometer was used t o measure temperature. T h e reactor was heated in a Lindberg Hevi-Duty, h o n e oven. T h e reac-
Time ( m i n u t e s )
6'x 1 / 8 ' S A Molecular s w a cclurnn nelum carrier - 50 ml/min
Time
Figure
(minutes)
6. Sample
Column temp,
chromatograms
70°C.,H W D ,
15OoC, 3 0 0 mA
tor temperature was controlled b y a Lindberg instrument console to within i l ° C in 1-2-in. beds. The reactor was mounted vertically resulting in convection gradients causing temperature variations in 12-in. beds as high as 20'C. This problem could be eliminated by careful adjustment of t h e outer zone heaters. However, t h e majority of runs were performed a t low contact times (0.07 t o 0.22 sec) corresponding t o catalyst depths of about 2 in. Thus, convection problems were minimal. The gas components of the feed stream were fed from pressurized tanks through a series of rotameters where the gases were blended to the desired concentrations. The exiting gas stream from the reactor contained sulfur vapor. To avoid plugging the lines downstream of the reactor with sulfur, the gas flowed through a line traced with heating tape to a n electrostatic precipitator where most of the sulfur was condensed arid removed. The precipitator consisted of a l l / r i n . diameter vertical aluminum tube, 24 in. in length, in which a copper wire was suspended centrally in the axial direction. The precipitator was powered by tapping the high-voltage source of a second-hand television set. T o ensure complete removal of elemental sulfur before reaching the analytical train, the gas stream was passed through a coil submerged in an ice bath. Figure 5 is a schematic of the experimental equipment. Analytical Equipment. Upstream and downstream gas compositions were analyzed with a Perkin-Elmer Ind. Eng. Chem Process Des. Develop., Vol. 12, No. 1, 1 9 7 3
13
0.05 t o 12 see. However, the reactions between CO and SO, -ool
TEMP*F
a
NOTE 1007 . reduction was redlzed
A
980 850 875 977 808 %4
0850 I 05
N
0.06
0.10
0.14 0.18 CONTACT TIME ( s e c o n d s )
0.22
Figure 7. Effects of contact time, temperature, and on percent sulfur dioxide reduction
CO WTIO
0
006
0 10
0 14 CONTACT TIME
0 18
CO
E
ratio
F 930 850 875 97 7 808 954 952
I 19 I32 I76
0 22
(seconds)
Figure 8. Effects of contact time, temperature, and CO ratio on percent carbonyl sulfide production
Model 900 gas chromatograph, containing a dual column system. CO,: COS, and SOa were analyzed on a 6-ft, X '/$-in. Poropak QS column, and S10 2 and , CO were analyzed on a 6-ft X '$-in. 5X Molecular Sieve column. Typical chromatograms are shown as Figure 6. Peak areas were determined with the use of a disc integrator. Experimental Procedure
Experimental runs indicated that COS and 0, reacted homogeneously to produce SO,(Querido, 1971). Therefore, to avoid the occurrence of this reaction, the oxygen must be removed prior to entering the main catalyst bed where CO and SO, will react. Furt,her experimentation showed complet,e conversion of Oe to CO,, using CO, was possible over CuO on alumina catalyst, a t temperatures greater than 840'F, and a t contact times greater than 0.189 see (Querido, 1971). Therefore, one technique for removing the oxygen would be to pass the stack gas through a "prebed" of CuO. K i t h the 02 removed, the stack gas entering the main cataCO, and SO*. It was this system lyst bed will contain COe, S,, that TTas investigated esperiment,ally for stack gas compositions correspoading to Case I3 in Table I as affected by the following variables: 1. CO ratio-0.63 to 2.09 2. Temperature-652-980OF 3. Contact time-Beinstock (1965) has suggested that contact t'imes be on the order of a few tenths of a second, which corresponds to space velocities of about l o 4 vol/vol:'hr. The experimental system was designed to allow contact t,imes from 14
Ind. Eng. Chem. Process Des. Develop., Vol. 12, No. 1, 1973
were observed to be extremely rapid, so that experimental times actually investigated were only 0.07 to 0.2 see. I n this paper, contact times are based on the interparticle void space measured a t the reaction conditions. 4. Catalysts-Ryason (1967) has suggested Group I B metals on alumina. Before attempting any experimental catalyst runs, the reactivity of the reactor and the silica alumina catalyst support were investigated and shown to be negligible. Initial studies involved the screening of the following catax '/$-in. cylindrical pellets): 10% CuO on lysts (all are alumina, 8% Cu on alumina, 475 Ag on alumina, 5y0MooBon silica-alumina ( 14-86y0), and the silica-alumina support itself. The Cu, CuO, and Ag catalysts rvere obtained from the Harshalv Chemical Co. and the silica-alumina support from the S o r t o n Co. The Moo8 catalyst was prepared in the laboratory by impregnating the support with an aqueous solution of ammonium molybdate. The catalyst runs were performed a t temperatures between 425" and 9i0°F, CO ratios ranging from 0.9-1.2, and a t contact times of about 0.4 see. The screening procedure indicated the following: Ag on alumina, although initially active, became readily poisoned with sulfur, as indicated by the presence of the metal sulfide. MOO:, catalyst also poisoned very quickly with sulfur. This catalyst is used in a modified form for the removal of SO,in the presence of H,, where H2S is formed and then reacts with SO2 to form elemental sulfur and water (hlea.de, 1969). CuO performed well only a t temperatures above l O O O O F with CO ratios greater thanunity. It was determined t h a t the excess CO probably reduced the CuO to the more active Cu form. Cu on alumina was very active, and was chosen for further study. Gas phase diffusion limitations were studied on the Cu on alumina catalyst during screening and were observed t o be virtually negligible (Querido, 1971). Results and Discussion
series of 34 runs was performed utilizing the Cu on alumina catalyst a t temperatures ranging from 652-980°F, a t contact times from 0.072-0.218 see, and a t CO ratios from 0.63 to 2.05. It' was generally observed that temperatures greater than 720°F were required t o realized 90y0 SO,reduction (Le., 90% of t'he initial SO, was reduced) even a t high CO rat'ios (2.05). Experimental SO, conversions and COS productions agreed very well with the thermodynamic equilibrium predictions for t'emperatures greater than 850°F and contact times greater than 0.19 see. At temperatures near 800°F, agreement between experiment'al data and thermodynamic predictions was generally within 15%. This indicated t'hat contact times greater than 0.2 see were required for t'he system to reach equilibrium. At 652'F and contact times on the order of 0.16 see, experimental results Tvere very far from equilibrium and SO2 reductions were less than 5 i % . Thus, in order for t h e reacting system to proceed with high rates (Le., short contact times) while achieving a minimum of 90y0 SOz reduction, temperatures above 800'F were required. Figures 7 through 9 present the results of the experimental series for temperatures greater than 800'F. I n Figure 7 , the percent SOz reduction (to both COS and sulfur) as a funct'ion of temperature and CO ratio is plotted against contact' time. At CO ratios greater than unit'y and a t A \
temperatures greater than 800"F, the reduction was very rapid, being completed a t contact times less than 0.22 see. As the temperature and/or CO ratio increased, 100% SO1 reduction was realized at shorter contact times. I n fact, a t CO ratios greater than 1.32 and temperatures greater than 800"F, complete reduction of SOz occurred within 0.09 sec. Figure 8 presents the experimental COS production (as percent of the inlet' SOz concentration) as a function of temperature and CO ratio. Generally, COS production increased with increasing CO ratio and temperature. At high CO ratios (1.76 and 1.82) the COS concentration increased with contact time, asymptotically approaching the maximum level predicted by thermodynamic equilibrium. A t CO ratios less than unity (0.803 and 0.850), the COS production appeared to be constant over the range of space times investigated, indicating t'hat the COS concentration attained steady state at contact times less than 0.0'7 sec. At CO ratios near unit'y (1.05 and 1.17) and a t relatively high temperatures (875" and 977"F, respectively), the COS concentration appeared to have attained its predicted thermodynamic equilibrium value between 0.15 and 0.20 see. Finally, the effect of temperature on the rate of COS production can be seen most readily by comparing the plots for CO ratios of 1.17 and 1.19. T h e former was run a t a temperature of 977°F resulting in a rapid approach to thermodynamic equilibrium. However, the latter run a t 808°F resulted in a lower rate of COS production and, hence, a slower approach to its equilibrium value. Figure 9 combines the results of the two previous plots and presents the effects of CO ratio and temperat'ure on the percent sulfur compounds remaining (unreacted SO2 plus COS produced as mole percent of initial SO?).T h e following conclusions can be drawn: -1total sulfur compound removal of greater than 90% was not realized at the invest'igated conditions. At high CO ratios (1.82, l . i 6 ) , 80-100~0of the SOz was ultimately converted to COS. At low CO ratios (0.803, 0.850) the percent remaining sulfur compounds appeared to be asymptotically approaching about 307, as the CO is exhausted. T h e minimum percent remaining sulfur compounds achieved (30'%) occurred a t CO ratios slightly above unity (1 .0&1. 19). Thus it appears that a iOy0 removal efficiency is feasible for contact times greater than 0.1 see and temperatures greater than 800°F. However, the CO ratio must be controlled quite closely (1.0 i 0.2) to maint,ain t,liis level of removal. The undesirable production of COS was the major problem encountered. For the process to meet the Public Health Service criteria (1967), the total sulfur compound removal efficiency must be great'er t'han 90%. Thus, COS had to be eliminated from the effluent or a t least its concentration had to be significantly reduced. Carbonyl sulfide is produced by the reaction between CO and Ss. This reactJion has been reported in the literature to be both homogeneous and heterogeneous (Chereponol- et al., 1966; Lewis and Lacey, 1915; Matlack et al., 1952; Parish, 1966; Pearson et al., 1932; Savin et al. 1968). *4t'temperatures above BOOOF, COS has been produced homogeneously in a tubular reactor (Lewis and Lacey, 1915; AIatlack et al., 1952; Pearsoii et al., 1932). Rj-ason (1967) has suggested that CuS may be the active sliecies for t,he reduction of SO2with CO. However, metal sulfides and alumina have also been reported to catalyze the reaction betweeii CO and S,producing COS a t temperatures below 600°F (Chereponov et al., 1966). Thus, Cu on alumiiia
1.17 1.B
0
0
01 0 06
I
I
0 10
I
0 18 (seconds)
0 14 CONTACT TIME
977
808
I
0 22
Figure 9. Effects of contact time, temperature, and CO ratio on percent remaining sulfur compounds
CuO on ALUMINA S ECO SPLIT N OPRY FLOW OR
(COS.SOe,
reactad i o S 2
co I
Cu on Aiumlna I
Figure 10. Three-stage sulfur dioxide removal process
catalyst appears to promote both reactions. Therefore, one method of eliminating COS would be to search for a catalyst that would be specific to the SOzreduction and would not promote the COS product'ion. However, in view of the possibile homogeneity of the COS reaction, an alternate met,hod of eliminating COS was developed. The method that proved to be successful utilized the reaction : 2COS
+ SO*=
3/2S2
+ 2c02
This reaction has been reported to be four times as fast as the CO-SO2 reaction over alumina catalyst a t I O O O O F , but a t substantially higher reactant concentrations than present in a stack gas (Lepsoe, 1940). However, by utilizing a more act'ive catalyst, such as Cu on alumina, the reaction proceeded rapidly even a t low reactant concentrations (1000 ppm). T h e SOs abat'ement process that was developed involved splitting the gas stream and separating the reactor into t'hree react'ors (hereafter designated as prebed, main, and secondary, respectively). X schematic of the process is shown as Figure 10. The entire gas stream would pass through a prebed of CuO m-here the oxygen is converted to COz. The exit stream would then be split, the major portion entering the main reactor containing Cu on alumina catalyst bed. Here, SO, would be reacted to sulfur and COS. The effluent from the main reactor along with the smaller split flow from the prebed would then pass into a secondary reactor also containing Ind. Eng. Chern. Process Des. Develop., Vol. 12, No. 1 , 1973
15
fn
a
40
-
30
-
20
-
F
\4 0
IO
'.O
I
0 a06
0.10
0-
- __ $_ - - o0- - -
I
I
I
0.14
0.18
0.22
-
8. 1.15 -0.495 R.0.85- 1.355
0.26
CONTACT T I M E
I
I
0.30
0.34
(aeconda)
Figure 1 1. Percent sulfur compounds remaining as function of contact time, CO ratio (R), beta, and temperature
t o
30t
TEMP: 583
- 603 * F
cos SATU 0.975 Io45 RdlO
I 006
Clor~litcotion:
R Z 0.9 R Z 1.0 1 . 0 ~ 0.9 ~ 7
I
0.10
I 014
LO HI MEO
A
1.015 0.856
0
1.08 1.14
0
I I I I 018 022 026 OX, CONTACT TIME (seconds)
I065 0.S60
I
034
Figure 12. Effects of COS ratio and CO ratio on percent remaining sulfur compounds
a Cu on alumina catalyst bed. Here, the COS formed in the main reactor would react with SO,to form sulfur. The gas stream entering the secondary reactor would contain COS, CO, and SO,.Thus, the following three reactions could occur:
Thermodynamically, all three reactions are favorable in the range of power plant boiler operating temperatures. Thus, a set of operating conditions where the kinetics of the second reaction would be much slower than those of t,he first or t,hird mas desirable. T o determine the operating temperature in the secondary reactor, a series of experimental runs over Cu on alumina catalyst was made with COS, CO, C o n , and SO,in the feed. Temperat'ures ranged from 986-506"F, CO ratios ranged from 0.55-1.35, and contact times ranged from 0.070.35 sec. The SO2 feed concentration was about' 1000 ppm for all runs. To facilitate the interpretation of the experimental data, t'he following parameters were defined: Relative feed ratio p 16 Ind.
=
COS concentration in feed CO concentration in feed
Eng. Chem. Process Des. Develop., Vol. 12,
No. 1, 1973
and, the COS ratio = Ncos/2Nso,, where &OS and X S O ~ represent the mole fractions of COS and SOz in the feed, respectively. Thus, the COS ratio specifies the amount of COS initially available to react with SOZ. The results of this series of runs are shown in Figure 11 which is a plot of percent remaining sulfur compounds (as mole yo of the initial SO2 and COS) as a function of contact time. In the range of contact times studied (0.07-0.35 sec), temperature had a marked effect on the direction and level of percent remaining sulfur compounds. At the relatively high temperature of 738"F, the remaining sulfur compounds increased with contact time indicating appreciable reaction between CO and S2yielding COS. At low temperatures about 500"F, the remaining sulfur c,ompounds decreased with increasing contact times which was indicative of very low rates of COS production and relatively slom-er rates of reaction of SO2 with CO and COS. At 590"F, no single curve could be drawn through the data points. Instead, the points were bounded by a region which exhibited a minimum of about' 9% remaining sulfur compounds, a t a contact time on the order of 0.18 sec. Figure 12 is a more detailed presentation of the data in this region. The data are presented in terms of the CO and COS ratios, where ratios below 0.9 were considered t o be low, above 1.0 high, and between 0.9 and 1.0 medium. I n general, high CO coupled with high COS ratios resulted in percent remaining sulfur compounds falling near the upper boundary of the region. Low CO and medium COS ratios formed the lower boundary. Referring back to the overall SOz removal process, the split stream entering the secondary reactor should contain CO concentrations corresponding to the initial CO ratio. The maximum sulfur compound removal efficiency in the main reactor of ?O% determined in the COSO, studies, indicated t'hat operation was restricted to CO ratios above unity. With the above restriction, from Figure 12 it is apparent that the minimum percent remaining sulfur compounds occurred a t low COS ratios. I n fact a minimum of about 13% (corresponding to a maximum sulfur compound removal efficiency of 87%) was realized a t low COS ratios (0.56-0.856) and a t contact times near 0.20 see. Thus, theoretically, it appeared feasible to achieve a total sulfur compound removal efficiency of greater than 90% for the overall process (Le., 7Oy0removal from t'he main reactor followed by 87% removal of the sulfur compound entering t'he secondary reactor). The preferential reactivity of the COS-SO? r e a ~ t ~ i oover n the CO-S2 react'ion a t temperatures below 600°K can be attributed to the sulfur equilibrium. .It temperatures below 600°K S g and Ss are the predominant sulfur species. Thus, the production of COS by the reaction of CO and S, is inhibited by the low concentration of S2a t these temperatures. The SO2 removal process was studied in the laboratory for a synthetic stack gas containing no 02,2000 ppm of SO:,1870 C 0 2 and CO ratios ranging from 1.03 to 1.25. With no 0 2 in the feed, onlj- two consecutive Cu-on-alumina-catalyst beds were considered in the experimental system. The first bed \vas contained in the furnace and reactor system described earlier. The secondary reactor was constructed from a 20-in. length of 1 x O . o G 3 - i ~315 stainless steel t'ubing, and was heated in a Lindberg, Hevi-Duty, single zone furnace. h series of 14 runs \vas made a t three different levels of CO ratios (1.03, 1,125, 1.25) tvith the main reactor maintained a t 900°F and the secondary reactor a t 600°F. The contact times ranged from 0.168-0.275 see in the main reactor and 0.198-0.201 sec in the secondary reactor. The parameter studied was the effect of the COS ratio in the secondary
reactor upon the efficiency of the process. This was easily accomplished by varying the split ratio defined as the flow rate to the main reactor divided by the flow bypassed to the secondary reactor, both measured a t the metering temperature. T h e results of these runs are presented in Figure 13. At a CO ratio of 1.03, sulfur compound removal efficiencies of greater t,han 90% were achieved over the ent'ire range of split ratios investigated. A maximum removal efficiency of about 9770 lyas realized at a split ratio of 5.5. The remaining two curves also exhibited masimums of greater than 90% removal. However, as the CO ratio increased, the masimums became sharper, decreased in magnitude, and were displaced toward smaller split ratios. The most significant feature of the esperimental results v a s that even a t the high CO ratio of 1.25, there esisted a split ratio which resulted in removal efficiencies greater than 90%. Thus, instead of requiring close control on the CO ratio direct'ly, the CO rat'io could be monitored upstream of the reactors and adjustments t o the split ratio would be employed to achieve the optimum remoyal efficiency. Controlling the split ratio would merely involve a valve adjustment whereas cont'rolliiig the CO ratio directly would be far more difficult, involving cont'rol of the excess air to the boiler. Sulfur is a notorious poison of metal catalysts. Therefore, to estimate the activity of the Ch 011 alumina catalyst, a 30-lir life test was performed on the dual reactor system. Sormally, signs of sulfur poisoning would occur in this time period. The test was performed at' a CO ratio of about 1.0 where high removal efficiencies would ensure a high level of sulfur in the reactors. The experimental results are shown in
LT
I
CO RATIO
K-
IO3 I 12
I
I
I
S K I T RATIO = (Main Reactor Flow Split to Secondary
]
mct~rln~
tamp.
Conclusions
From the foregoing discussion of experimental results, the following can be concluded: Complete reduction of SO2 with CO was achieved over Cu on alumina catalyst a t temperatures greater than 800"F, CO ratios above unity, and at contact times on the order of 0.2 sec. However, a deleterious side reaction between elemental sulfur and CO to form COS also occurred. This side reaction limited the total sulfur compound removal efficiency to about 'io%, where total refers to COS and unreacted SO*. Results of the catalyst screening study indicated that from the series of CuO on alumina, Cu on alumina, X g on alumina, N o 0 3 on silica alumina, and the silica alumina catalyst support itself, only Cu exhibited sustained activity toward the SO2-Co reaction. The esperimental SO* conversions and COS productions agreed very \vel1 with the thermodynamic equilibrium predictions for temperatures greater than 850'F and contact times greater than 0.19 sec. It has been shown that greater removal efficiencies can be achieved with the employment of the dual Cu on alumina catalyst reactor system. By adjustment of the split f l o ~to the secondary reactor, maximum sulfur compound removal efficiencies greater than 90% were possible for CO ratios as high as 1.25. A maximum sulfur compound removal efficiency of 9i% was realized a t a CO ratio of 1.03 Jvith the dual reactor system. It was observed that as the CO ratio was increased, the maximums became sharper, decreased in magnitude, and were displaced toward smaller split ratios. However, even a t the relatively high CO ratio of 1.25, a maximum efficieiicy of 92.67, n a s relaized. After the operating conditions for the dual reactor system had been established, the effect of time on conversion was studied. The Cu on alumina catalyst indicated no appreciable activity decay during a 30-hr run.
Literature Cited
Split R a t i o
Figure 13. Effect of split ratio on percent sulfur compounds removed
n
MAIN
3
Figure 14. The sulfur compound removal efficiency reached its steady-state value of about 95.1% after 4 hr and maintained that level for the duration of the run. Hence, the catalyst did not show any appreciable decay for the 30-hr period.
SECONDARY
C O R a t i o : 0,997 Split R a h o : 6.1
Avg % Sulfur Oompound A m o v e d : 95.10
30 Tlme (hours)
Figure 14. Copper on alumina catalyst life test
Baker, R. A., Doerr, R.C. Ind. Eng. Chem. Process Des. Develop., 4 (2) (April 1965). Beinstock, D., Field, J. H., Katell, S., Plants, K. D., A P C A J., 15 (10) 459 (1965). Berkowitz, J., Chupka, W. A., J . Chem. Phys., 40 (2), p 287 (January 1.5, 1964). Berkowitz, J., Marquart, J. E., ibid., 39 (2), p 275 (July 15, 1963). Patent 181,064 (April 15, Chereponov, S. I., et al., U.S.S.R. 1966). Ferguson, J. B., J . Amer. Chem. SOC.,40, 1626 (1918). Gamson, B. W., Elkins, R. H., Chem. Eng. Progr., 49 (4), 206 (1953). Hangebrauck, R. P., Spaite, P. W A P C A J . , 18 (1) 5 (1968). Journal of Thermochemical (JA-TAF) Tables, Dow Chemical Co., Distributed by the Clearing House for Federal, Scientific and Technical Information, U.S. Dept. of Commerce, Natl. Bur Stand., Institute for Applied Technology (Code 410.14), Springfield, Va. Lepsoe, R., Ind. Eng. Chem., 32,910 (August 1940). Lewis, G. N., Lacey, W. K.,J . Amer. Chem. Soc., 37, 1976 (1915). Matlack. G. h l . . et al.. Proc. Iowa Acad. Sci.. 59.202 (1952). Meade, C. W., Private Communication, 1969. Xleyer, B., Ed., "Elemental Sulfur," p 127, Interscience, Kew York, N.Y., 1965. Ind. Eng. Chem. Process Des. Develop., Vol. 12, No. 1 , 1973
17
Munro, A. J. E., Madsin, E. G., Brit. Chem. Eng., 12 (3), 364 (1967). Parish, G. S. Patent 3,416,893 (1966). Pearson, T. G. et al., J . Chem. SOC.,1932,p 660. Perry, R. H., bhilton, C. H., Kirkpatrick, S. D., Eds., “Chemical Engineering Handbook,” RlcGraw-Hill, Kew York, N.Y.,
U.S. Department of Health Education and Welfare, “Air Quality Criteria for Sulfur Oxides,” Washington, D.C., 1967. lJ.S. Department of Health Education and Welfare, “Control Techniques for Sulfur Oxide Air Pollutants,” Washington, D.C. 1968. White, W.S., et al., J . Chem. Phys., 28 ( 5 ) ,751 (1958).
QuL;Y&, R., PhD Thesis, University of hIassachusetts, 1971. Ryason, P. It., Harkins, J., J . APCA, 17 (12), 796 (1967). Savin, V. P., et al., (U.S.S.R.), Khim. Prom., 44 (7), 550 (1968). Saxtaxtinskij, G., Bajand-Guliev, A. I., Azer. Khim. Zh. (6), 123 (1965). Slhck, A. V., Chem. Eng. (December 4, 1967).
RECEIVED for review July 29, 1971 ACCEPTED September 8, 1972
1a m
Work supported by Contract PHS-AP00791 with the Public Health Service, Department of Health, Education and Welfare.
Design of Closed-Circuit Grinding System with Tube Mill and Nonideal Classifier Masaaki Furuyafl Yoji Nakajimaf2and Tatsuo Tanaka Department of Chemical Process Engineering, Hokkaido Cniversity, Sapporo, Japan
Based on the comminution kinetics and material balances, the finished product size distribution of a typical closed-circuit grinding system with a tube mill and a nonideal classifier are calculated for a number of system parameters. The calculated results are summarized as simply and generally as possible with the aid of analytical consideration. Some simplified correlations are presented in a few figures which may be useful for designing new system or for improving old system, if data on mill and classifier performances are available. In connection to it, a simple method for determining the parameters of the commiwtion kinetics from open-circuit milling data i s suggested.
T h e authors (1971) reported an analytical discussion on closed-circuit grinding system with a n ideal classifier. As a matter of fact, however, a n ideal or clean-cut separation of particles cannot be expected to occur in any industrial classifier, and therefore the system analysis described in the previous paper is not always applicable to design purposes. I n this paper, a certain improvement is made in the comminution kinetics and the fractional recovery curves of industrial classifiers are represented in a simple function having several parameters. The finished product size distributions of a typical closed-circuit grinding system are numerically calculated for a number of combinations of the system parameters such as circulating load, sharpness of classification, and so on. However, the main purpose of this work does not consist in a n accurate prediction of the performance of a given closedcircuit grinding system, but consists in a general extension of the comminution kinetics for solving the system design problem; the accuracy of the prediction may be sacrificed to some extent for this purpose. I n this paper, the calculated results are summarized as generally and simply as possible with the aid of analytical consideration, and some simplified correlations represented in a few figures are presented which may be useful for designing a new system or improving the old system. 1
Present address, Shin-etsu Chemical Co., Gumma, Japan.
* T o whom correapondence shoi11d be addressed.
18 Ind. Eng. Chem. Process Des. Develop., Vol. 12, No. 1 , 1973
Mathematical Preparation
For the steady state of the operation of the grinding system shown in Figure 1, the following equations are derived from material balances :
+ + + CLfTII(1 + = (1 + C d ( 1 - avo
fE = (fF
fD = (fP fP
CLfT)/O
CL)
(1)
CL)
(2)
(3)
where f denotes the size frequency distribution by weight and the subscripts, D, E , F , P , and T , respectively, refer to the mill discharge, mill entry, make-up feed, finished product, and tailings; CL is the circulating load; and 7 is the fractional recovery of the classifier as a function of particle size. As suggested in the previous paper (Furuya et al., 1971), the flow mechanism in a n industrial tube mill may reasonably be assumed as a plug flow so far as the product size distribution is in question. Consequently, the batch milling equation (Austin and Klimpel, 1964), Equation 4, is directly usable to relate fE and jD without any consideration on the residence time distribution of particles in the mill:
where R(y,t) is the cumulative oversize by weight a t the size y and the milling time t , x m the maximum size in feed material,
