Dual Bed Catalyst for Simultaneous Reduction of Sulfur Dioxide and Nitric Oxide with Carbon Monoxide Ajay Sood* and J. R. Kittrell Department of Chemical Engineering, University of Massachusetts. Amherst, Massachusetts 01002
For the simultaneous reduction of SO? and NO by CO, a dual-bed catalyst is discussed, comprised of iron-chromia as a first bed, followed by activated alumina as a second bed. Undesirable COS formed in the first bed reacts with SO2 in the second bed to yield sulfur. Greater than 90% reduction of both SOz and NO at -700°F and high space velocity (-20,000 h r - l ) is demonstrated with tubular flow reactors using simulated flue gases. Moreover, variations in catalyst temperature can adjust the COS:SO2 ratio to that required stoichiometrically and also counterbalance any changes due to fluctuating CO feed level.
Sulfur dioxide and nitric oxide emissions from stationary sources represent a significant fraction of the total emission of these air pollutants in the U. S. A. A number of dry and wet control processes have been proposed. These require absorption, adsorption, or reaction of the pollutants with a liquid or solid which must be subsequently regenerated or disposed of properly; moreover, several chemical processing steps are necessary before the sulfur value can be recovered, frequently in a form (such as dilute sulfuric acid) which is not easily marketable or storable. A dry process based on the reduction of sulfur dioxide and nitric oxide in stack gases with carbon monoxide over a copper-alumina catalyst has recently been described by Quinlan and Kittrell (1973). This process has the advantage of not requiring regeneration of reactants and producing elemental sulfur which can be easily stored and marketed. Duffy (1972) has compared the economic feasibility of this process to other first generation processes being currently investigated and concluded that the CO reduction process is attractive if a suitable catalyst system can be developed. Important reactions occurring during the simultaneous reduction of sulfur dioxide and nitric oxide by carbon monoxide are Co 2CO
+
NO = CO2
+ so2
+
= 2c02
Y2N2
(1)
+ ‘/zsz
(2)
co + %s2 = cos (3) The first two reactions are the reduction of NO and SO2 to N2 and elemental sulfur, respectively. Equation 3 represents the formation of COS by an undesirable side reaction between CO and the elemental sulfur produced in eq 2. Ryason and Harkins (1967) observed some COS formation while studying the catalytic reduction of SO2 with CO over several catalysts. Khalafalla and Haas (1972) similarly noted COS production with a bifunctional ironalumina catalyst. Querido and Short (1973) also observed the deleterious side reaction between elemental sulfur and CO to form COS in their study of a commercial copperalumina catalyst. Quinlan, et al. (1973a1, further defined the relationship between COS production and SO2 conversions for this commercial copper-alumina catalyst and demonstrated that a maximum of 80% sulfur removal was possible. Goetz, et al. (1974), screened a large number of laboratory prepared and commercial catalysts for the simultaneous reduction of SO2 and NO with CO and found 180
Ind. Eng. Chem., Prod. Res. Develop., Vol. 13, No. 3, 1974
that COS was produced by all the catalysts tested. Hence, it appears that all catalysts reported in the literature, although differing in activity and selectivity, have a common problem-the production of undesirable COS. If the CO reduction process is to be viable, it is necessary to find a method of eliminating the COS formed. One method, suggested by Querido and Short (1973), involved using two reactors in series containing the copper-alumina catalyst. The feed containing CO and SO2 was split between the main and secondary reactor. The COS produced in the main reactor was then further reacted with SO2 in the secondary reactor according to the reaction 2cos
+ soz = V2sZ4 2
~
0
~
However, since CO and SO2 were present in the feed to the secondary reactor, some COS was formed simultaneously according to eq 2 and 3 on the copper-alumina catalyst. Because of this COS formation in the secondary reactor, Querido and Short (1973) found that the IeveI of CO in the feed had to be controlled very carefully. The CO level should not be more than 3% above stoichiometric requirements in order to obtain higher than 90% total sulfur removal (based on COS S02). At higher CO levels, a high sulfur removal rate was possible only if the ratio of the feeds to the main and secondary reactors was precisely controlled. T o overcome these problems, it was obvious that a catalyst which did not promote the CO-SO2 reaction but was highly active for the COS-SO2 reaction would be desirable for the secondary bed. As pointed out by Khalafalla and Haas (1972), the CO adsorption occurs on the metal in the catalyst; in the absence of the metal no CO adsorption or CO-SO2 reaction can occur (this was verified experimentally in our laboratory). Numerous studies have evaluated catalysts for the reaction of COS and S O Z , but the primary metal-free catalysts reported have been activated alumina and bauxite (Agren and Lindhe, 1935; Ferrante, 1946; Gamson and Elkins, 1953; Hurlburt and Davis, 1970; Lepsoe and Mills, 1937; Roesner, 1937; Sefton, 1971; West and Conroy, 1965). After an extensive catalyst screening program, a high surface area (380 m2/g) activated Kaiser alumina, KA201, was found to be a very active catalyst for the COSSO2 reaction. Since it had no metal component, it did not lead to any COS formation (according to eq 2 and 31, but only to COS reduction (according to eq 4) a t the CO levels important in stack gas SO2 abatement. Thus, an alternate method of eliminating the COS problem based on a dual-bed catalyst system is proposed
+
in this paper. A single reactor is used with downward flow of the simulated stack gases through two successive beds of catalysts placed in the reactor. The first catalyst bed consists of a commercial Girdler iron-chromia catalyst, G-3A, which was found to be a highly active catalyst for the reactions of eq 1, 2, and 3 by Goetz, et al. (1974). This will henceforth be referred to as the iron catalyst for convenience. The second bed consists of the Kaiser alumina and promotes the reaction shown in eq 4. Here, no splitting of the feed stream is necessary, so that the entire effluent of the first bed passes through the second bed. It is shown how variation of the catalyst temperature in this dual bed system can be used to adjust the ratio of COS t o ' SO2 entering the second bed as well as to compensate for any changes in the CO concentration of the feed.
Experimental Section The experimental system was designed so that six parallel tubular reactors could be operated simultaneously a t a common temperature with an identical feed gas composition; the upstream and downstream gas concentrations could be analyzed without disturbing the system. Thus, six different experiments could be performed simultaneously on this equipment. The equipment included a blending manifold, six reactors contained in a common furnace, room temperature and ice bath precipitators, a sampling manifold, and a gas analysis system discussed earlier (Goetz, et al., 1974). In the blending manifold, a gas mixture of the desired composition was obtained from pure, compressed, dry gas cylinders of N2, NO, CO, and S02; no water was added in any of these experiments. Type 304 stainless steel tubing and valves were used prior to the analytical train. The gases were blended through micrometering valves and rotometers and the resultant mixture split into seven streams-one line for determining the upstream gas concentration and the others for feeding the six reactors. The flow through each of these seven lines was regulated by micrometering valves, rotometers and pressure gauges. The upstream pressure in all the reactors was approximately 250 mm Hg. As a safety precaution, the entire equipment was enclosed in a hood and vented to an exhaust fan. For the six reactors, titanium was used as the material of construction since Quinlan, et al. (1973b), had reported that it does not promote the reductionof SO2 or NO by CO. The reactors were made from lis-in. I.P.S. titanium pipes having a n internal diameter of approximately in. Each reactor had a 40 mesh stainless steel screen in the middle to support the catalyst and a union coupling a t the top for charging or removing the catalyst. The average particle size of both the iron and alumina catalysts tested was chosen to be 20/30 mesh. This gave a particle to reactor internal diameter ratio of 1:9, which was slightly lower than the minimum recommended by Dowden and Bridger (1957) in order to ensure plug flow through the catalyst bed. For the experimental conditions used in this study, theoretical calculations as well as experimental results indicated that internal and external film diffusion limitations were negligible (Sood, 1974). All the reactors were vertically supported in the same Lindberg Heavi-Duty, three-zone, hinged tube furnace. The furnace was used in conjunction with a Lindberg Heavi-Duty three-zone temperature controller so that the furnace temperature could be set to any desired level. The controller settings were adjusted to operate the catalyst bed isothermally within *3"F. The actual catalyst temperature was measured by monitoring the output of three chromel-alumel thermocouples, mounted on the outer
wall of each titanium reactor, with a Leeds and Northrup potentiometer. The effluent from the reactors contained elemental sulfur which had to be removed to prevent plugging of the downstream lines and interference with the chromatographic analysis. Hence the reactor effluents were passed through room temperature precipitators and an ice bath to condense and remove the elemental sulfur before going to the sampling manifold. This manifold was designed to permit sampling of either the upstream or any one of the reactor downstream lines without creating any disturbance in the system. The analysis for nitric oxide was performed by Dynasciences Pollution Monitor NS130 equipped with a SO2 scrubber. This instrument was calibrated with a standard nitric oxide-nitrogen mixture. The analysis for CO, COS, and SO2 was done with a Varian Aerograph Model 1860-30 chromatograph, using helium as the carrier gas, a detector temperature of 115"C, and a column temperature of 68°C. The CO was analyzed with a Molecular Sieve 5A column, while COS and SO2 were analyzed with a Porapak QS column. Both columns were 6 ft X Yg in. and employed 80/100 mesh packing.
Thermodynamic Equilibrium Predictions Thermodynamic predictions of equilibrium compositions to be expected from a variety of feed gas compositions were made a t several temperatures of interest and 1 atm total pressure. These calculations were based on a free energy minimization technique which utilizes a modified steepest descent search method as proposed by White, et al. (1958). The chief advantage of this technique is that the actual chemical reactions occurring need not be specified but only the reactants and products involved. All free energy data were taken from the JANAF tables. Separate computations for the two catalyst beds were necessary since the reactants for the two catalysts were quite different from each other. For the first bed, with a feed containing CO, S o n , NO, and N2, the equilibrium results and conclusions are identical with those reported by Quinlan, et al. (1973b), and hence will not be repeated here. For the second bed. a COS ratio is defined for ease of discussion as COS ratio =
COS i n feed ( p p m ) 2[S02 in feed (ppm)]
(5)
A COS ratio of unity indicates that the COS and SO2 in the feed are in stoichiometric proportions, as required by equation 4. Equilibrium computations for a second bed feed containing 2050 ppm of COS plus S02, 3350 ppm of C02 and the balance N2 were performed. Shown in Figure 1 are the COS and SO2 remaining a t equilibrium a t various COS ratios. It is evident that the reaction goes to completion a t a COS ratio of unity, while above and below this stoichiometric requirement, SO2 and COS are the limiting reactants, respectively. For this calculation, only S2 and SS forms of sulfur were considered, as explained elsewhere (Querido and Short, 1973). The curves indicate that a minima in the total sulfur compounds remaining exists a t a COS ratio of unity. Lowering the temperature sharpens this minima, but the equilibrium conversion is always higher than for the Claus reaction. Thus, the COS ratio of the feed to the second bed as well as the catalyst temperature can be expected to have a very significant effect on the percentage sulfur compounds remaining in the effluent. This prediction was substantiated by experiments reported here. The thermodynamic calculations did not predict any Ind. Eng. Chem., Prod. Res. Develop., Vol. 13, No. 3, 1974
181
1000
-
903
-
/
I
6o
/
t
/ / 980'F
/
0
/
$-
260'-620'F
800 -
40
a I
/
w
1
700 -
/
30
In
n
z
3 0
. 600 f
5
5
so0
20
2
0
8
-
LT
10
U 3 3 ' " 0
N400-
0
300
-
200
-
103
-
#
0
cos
1
COS
I
I
2
3 N
Figure 1 . Dependence of equilibrium behavior of 2COS + S O P = 2C02 + 3 / ~ on S ~process conditions. Feed contains: COS + SOP = 2050 ppm, COz = 3350 ppm, N2 balance.
100
-
w I-
formation of CS2 and CO as products a t temperatures between 260 and 800°F and COS ratio less than unity, thus negating the possibility of the reactions -c
2cos
=
2 -s,
i
90
-
cs2
(6)
2co
(7)
0
zz z
70-
z
W
60 -
N v) 0
At higher COS ratios, formation of CO and CS2 in trace quantities only was predicted. No formation of CO or CS2 was observed in the experiments with Kaiser alumina. From Figure 1, it can be seen that greater than 96% total sulfur compounds removal is possible thermodynamically a t a COS ratio of unity and temperatures below 980°F. There is, of course, an effect of COZ level on equilibrium sulfur compound removal, as expected from eq 4 and shown in Figure 2 . As demonstrated by the results to follow, sulfur removal higher than 95% was achieved experimentally in this dual bed system at high space velocities.
Results A series of experiments was performed using dry cylinder gases a t space velocities (based on gas reaction conditions) from 20,000 to 100,000 hr-l, with a nominal feed composition of 2000 ppm of SO2, 600 ppm of NO, and CO levels from 5700 to 6500 ppm (about 1.3 times the stoichiometric CO required). The gas flow rate in all the experiments was about 950 standard cc/min. No N20 in the reactor effluents could be detected, indicating that the NO was reduced to N2. Also thermal decomposition was negligible. Furthermore, since the NO conversion to Nz exceeded 950/0 for all these experiments, the discussion here will emphasize only SO2 and COS levels in the reactor effluents. The presence of N O was not found to promote the CO-SO2 reaction. First, the activity of the Kaiser alumina alone for promoting the reduction of SO2 and NO with CO was checked a t a temperature of 665°F and a space velocity of 37,000 hr-l. No conversion of any of the reactants was observed, nor was any formation of carbonyl sulfide detected. This confirmed that a metal component in the cataInd. Eng. Chern., Prod. Res.
1400
-
+
50
REACTOR 2
v,
182
1200
'F
P $ 80
coz
1000
a U 0
=
800
Figure 2. Dependence of equilibrium sulfur compounds (COS + SOP)remaining upon COz level. Feed contains: COS = 1367 ppm, SO2 = 683 ppm, COz as shown, NZbalance.
--
RATIO
2cos
600
TEMPERATURE,
E l
0 U
400
ZOO
Develop., Vol. 13, No. 3, 1974
0
A
a
40
39 I R O N - C M R O M I A 2 N D B E D l g ALUMINA I ST BED
A
t-
c 30
-
20
-
.A
I
Bc
.
A '
-___-_---~
b.
/
A
I
I
I
I
I
lyst was necessary for CO adsorption and that the alumina by itself cannot promote the CO-SO2-NO reaction under the experimental conditions chosen. Next, a series of experiments was designed to determine the relative proportions of the catalysts necessary in the two beds for complete removal of the total sulfur compounds (COS and SO2) in the reactor effluent. The battery of six tubular reactors, operating in parallel, and having an identical feed gas composition and almost the same temperature was extremely useful in conducting this set of experiments. The six reactors were charged with different proportions of the two catalysts and the test was conducted a t a CO level approximately 1.3 times the stoichiometric requirements, a feed gas flow rate of 950 cc/ min (at 70°F and 760 mm Hg pressure) in each reactor and several catalyst temperatures ranging from 635 to 694°F. In Figure 3, the total sulfur compounds (COS + SO21 remaining, expressed as a percentage of SO2 in the feed, are plotted against the catalyst temperature for reactors 1 and 2. Reactor 1 was charged with 1 gpf alumina above 3 g of iron catalyst (this has already been shown to be equiva-
901
REACTOR
COS
I
eo -
O
SO2
-
.
loo
0
0
0
a
0
1)
REACTOR 3 N
s
70 N 0
IW
5
60-
0
Y
50
z. 7 0 -
SEE F I G 3
-
W
0
z -
3
2
A
60-
REACTOR 4 I S 1 BED 19 IRON-CHROMIA
(r
30
N
-
\ \
W
0 v)
0 50-
\
\ \
2ND BED 3 g ALUMINA
v)
+
v)
v)
0
40-
20 -
c
A
z
W z
i:
\
5 40-
N
LT
00-
Y
I-
z
IS1 BED 3 Q ALUMINA 2ND BED I 9 IRON-CHROMIA
90-
10
?
-
30-
LL W
..
0 -
~~
630
640
650
CATALYST 630
640
650
660
670
680
690
7a3
CATALYST T E M P E R A T U R E , 'F
Figure 4. COS and SO2 remaining in effluents of reactors 1 and 2.
lent to 3 g of iron alone). Reactor 2 was charged with 1 g of alumina below 3 g of iron catalyst. Except for the relative location of the two catalyst layers, all other test conditions were identical. The space velocity for the combined bed was 24,000 hr-1, while for the iron and alumina beds separately it was 36,000 and 75,000 hr-l, respectively. From the plot, it is apparent that the two reactors differ markedly in their effluent sulfur compound levelsreactor 2 achieves about 70% conversion of the entering SO2 whereas reactor 1achieves 25% conversion a t best. In Figure 4, the effluent COS and SO2 levels for reactors 1 and 2 are shown by solid and broken lines, respectively. It is to be remembered that the effluent composition of reactor 1 approximates the concentration of gases entering the lower alumina layer of reactor 2. Examining reactor 1 effluent compositions, we note that the proportion of COS to SO2 is much higher than the 2:l ratio required by the stoichiometry of eq 4. Hence we would expect that reactor 2 should have significantly lower effluent levels of COS and SO2 than reactor 1 but be unable to convert all the COS entering the alumina bed for lack of adequate SO2 being available. As shown by the broken lines in Figure 4, this is indeed the case if the temperature of the alumina bed is higher than 650°F. Furthermore, since the COS production is limited by the iron bed of reactor 2, higher conversions than shown in Figure 3 should be possible. One way to decrease the quantity of COS formed relative to the SO2 remaining is to decrease the amount of iron catalyst charged to the reactors. Accordingly, reactor 3 was charged with 3 g of alumina above 1 g of iron catalyst (equivalent to 1 g of iron alone, as explained earlier) and reactor 4 was charged with 3 g of alumina below 1 g of the iron catalyst. For these reactors a 20,000 hr-1 space velocity on the total bed was equivalent to a 103,000 hr-' space velocity on the iron catalyst. The results are illustrated in Figure 5. N o sulfur formation is achieved in reactor 3 even though some SO2 is converted to COS (Figure 6). In contrast, reactor 4 shows approximately 60% sulfur removal a t 686°F. Figure 6 shows the effluent compositions of reactors 3 and 4. Insufficient COS is formed
660
670
TEMPERATURE
600
690
70)
, O F
Figure 5. Total COS + SO2 remaining in effluents of reactors 3 and 4. Test conditions: combined S.V. 20,000 hr-l; feed contains 6000 ppm of CO, 600 ppm of NO, 2000 ppm of SOz, balance N P .
eo
-
70
-
60
-
N
0 v)
!W _J
z
\
50-
LL
0
8. N
0 U
30v)
0 0
c
20 -
W 3
6%
1
I
I
640
650
660
I
I 670
CATALYST TEMPERATURE
680
I 690
, O F
Figure 6. COS and SO2 remaining in effluents of reactors 3 and 4.
relative to the SO2 remaining in the reactor 3 effluent stream, with the result that the alumina of reactor 4 reacts all this COS and there is an excess of SO2 remaining, thus limiting the sulfur conversion of reactor 4 to 60%. The temperature of reactor 4 could be increased to about 710°F to raise the sulfur conversion to the 90% level but this was not attempted here. In order to increase the COS formation, the amount of iron catalyst was increased to 2 g. Thus, reactor 5 contained 2 g of alumina above 2 g of iron (equivalent to 2 g of iron alone) and reactor 6 had 2 g of alumina below 2 g of iron catalyst. The space velocity for the combined bed was 22,000 hr-l, while for the iron and alumina beds separately it was 54,000 and 36,000 hr-l, respectively. The results are summarized in Figure 7 . For reactor 5 , the highest Ind. Eng. Chem., Prod. Res. Develop., Vol. 13, No. 3, 1974
183
100
~
REACTOR
~
~
CO L E V E L ,ppm
90
_
_
_
_
_
_
~
A
COS
SO2
REACTOR 5
N 80
I ST BED
2 Q ALUMINA
2ND BED
2
8 IRON-CIIROMIA
6300
70 0
0 v)
+ W
A
70 A 1
5
k
?
60
a-'
U
0
\
a-'
\
u 50
z
\
I
\
5r
\
40
v)
6
REACTOR
\
N
0
I S T BED
\
m
0
I A
Y)
1
0
2 9 IRON-CHROYIA
ZNDBEO 29
\
u
\
-J
; 4
\
10
I
640
650
I 660
A-0
I 670
TEMPERATURE,
IO
U
W
/
-8
'-* I
3
I
\
CATALYST
w z
I 1
\
A A I 6eO
A
L
I
\
v, 0 20
A 0
I
A '
m
30
I I
ALUMINA
'A
30
+
0
40
N
I
I
REACTOR
B
1
/'
I
I
700
710
I 6W
F'
+
Figure 7. Total COS SO2 remaining in effluents of reactors 5 and 6. Test conditions: combined S.V. 22,000 hr-1; feed contains 6000 ppm of CO, 600 ppm of NO, 2000 ppm of SOZ,balance Nz. 60
80
70 0
N
60
0
I-
w
N
-z 50
v) 0
20
U
0
0
8
v)
40
8
N
10
c
0 v,
w
r 30
$ In
0
W
020
c
z
660
w
6
lo
0 I
0
I
I
I
I
640 650 660 670 680 C A T A L Y S T T E M P E R A T U R E , 'F
I
690
I 700
Figure 8. COS and ,302 remaining in effluents of reactors 5 and 6.
sulfur conversion obtained was 25% a t 690°F. If the temperature had been increased further for this reactor, the remaining sulfur compounds would have first decreased and then increased due to excessive COS formation. The per cent remaining sulfur compounds for reactor 6 exhibit a minima a t a temperature of about 677°F. At this temperature, five data points were taken with sulfur compounds removal ranging from 92 to 99% of the feed SO2 level. This high sulfur removal was achieved by obtaining the proper ratio of COS to SO2 in the effluent of the iron bed by adjustment of the catalyst temperature, as shown in detail in Figure 8. If the temperature of reactor 6 is less than about 677"F, the conversion to sulfur decreases be184
670 CATALYST
Ind. Eng. Chern., Prod. Res.
Develop., Vol. 13, No. 3,1974
680
690
TEMPERATURE
,OF
Figure 10. COS and SO2 remaining in effluent of reactor El. Test conditions: first bed, 2 g of iron-chromia catalyst; S.V. 54,000 h r - l ; second bed, 2 g of alumina catalyst; S.V. 36,000 hrr1. Feed contains CO as shown, 640 ppm of NO, 2100 ppm of SOz, balance Nz.
cause of remaining SO2 and if the temperature is above this level, the conversion decreases due to remaining COS. The results demonstrate that the catalyst temperature (and hence the COS to SO2 ratio) must be carefully controlled to achieve greater than 95% removal of the total sulfur compounds. The effect of raising the inlet CO concentration is to increase both the rates of SO2 conversion to sulfur and also sulfur conversion to COS. This is because the concentration of one of the reactants, CO, in the reactions shown by eq 2 and 3 is being increased. However, in theory, changes in the CO concentration of the flue gas should cause no problem in the dual bed catalyst system, for changes in catalyst temperature should make any required adjustment of COS to SO2 ratio entering the alumina bed,
thereby maintaining a t least 90% conversion of SO2 to sulfur. Several experiments were performed to confirm this expectation. These runs were made a t the same nominal conditions as used for Figures 7 and 8 except that the CO level was varied between 1.1 and 1.3 times the stoichiometric requirement. The results are presented as the percentage COS and SO2 remaining in the effluent (based on SO2 in the feed), so that sum of these two represents the total unconverted sulfur compounds. These values are plotted against the catalyst temperature in Figures 9 and 10. In Figure 9, the effluent concentrations of COS and SO2 are shown for reactor A, which contained 2 g of iron catalyst only. Note that at all the temperatures and CO levels, the COS and SO2 in the reactor effluent was roughly present in the required 2 : l stoichiometric ratio. In Figure 10, the results for reactor B, which was charged with 2 g of alumina below the iron catalyst are shown. These results demonstrate that conversion of SO2 to sulfur can be achieved a t the 90% level, with COS being preferentially driven to zero, if desired. Note that the 6300 ppm of CO points have a high COS level; it has already been shown that if the catalyst temperature is lowered, this COS level can be driven to zero and hence that experiment was not repeated here. Conclusions
These experiments have clearly demonstrated the feasibility and usefulness of the dual bed catalyst concept. The results have shown that it is possible to operate a single reactor with two catalyst layers a t temperatures less than 700"F, high space velocities (at least 20,000 hr-1 on the combined catalyst bed), and with greater than 90% conversion of both SO2 to sulfur and N O to N 2 . Specifically, these results show that the ratio of COS to SO2 entering the second alumina bed can be adjusted to
the desired stoichiometric ratio of eq 4 by changing the catalyst bed temperature. If the effluent SO2 concentration is too high, the reactor temperature is raised. On the other hand, if the effluent COS concentration is too high, the reactor temperature is lowered. Furthermore, it has been shown that relatively minor changes in catalyst temperature are sufficient for controlling wide ranges of entering CO concentration. L i t e r a t u r e Cited Agren, P. H. W., Lindhe, J. H. 0.. Swedish Patent 84,978 (Dec 3, 1935). Dowden, D. A.. Bridger, G. W., Advan. Catal.. 9, 699 (1957). Duffy, R. T., M.S. Thesis, University of Massachusetts, 1972. Ferrante, C., Italian Patent 412,812 (Feb 16, 1946). Gamson. B. W., Elkins, R. W., Chem. Eng. Progr., 49 ( 4 ) , 203 (1953) Goetz, V . , Sood, A., Kittrell, J. R.. Ind. Eng. Chem., Prod. Res. Develop., 13, 110 (1974). Hurlburt, ti. Z., Davis, M . H.. Jr., German Patent 2,019,660 (Nov 5, 1970) KhaiafHila, S.E., Haas. L.A,, J. Catal., 24, 121 (1972). Lepsoe, R . , Mills, J. R., U. S. Patent 2,080,360 (May 11, 1937). Querido. R., Short. W. L., fnd. Eng. Chem., Process Des. Develop., 1 2 , 10 (1973). Quinlan, C. W., Kittrell. J. R.. paper presented at the 39th Annual Winter Symposium of the American Chemical Society, Chicago, Ill,, Jan 18, 1973. Quinlan, C. W.. Okay, V. C., Kittrell, J. R., lnd. Eng. Chem., Process Des. Develop., 12, 107 (1973a) QUinlan, C. W., Okay. V. C.,Kittrell, J. R., fnd. Eng. Chem., Process Des. Develop., 12, 359 (1973b) Roesner, G.. Chem. Fabriic, 101 (1937) Ryason, P. T . , Harkins, J., J. A i r Pollut. Contr. Ass.. 17, 796 (1967). Sefton, V . B., U. S. Patent 3,579,302 (May 18, 1971). Sood, Ajay, Ph.D. Thesis, University of Massachusetts, 1974. West, J. R., Conroy. E. H., U. S.Patent 3,199,955 (Aug 10, 1965) White. W. B., et a/., J . Chem. Phys.. 28 (5), 751 (1958). Receivedfor reuieu: F e b r u a r y 4, 1974 Accepted J u n e 3,1974 T h i s project has been financed in p a r t w i t h Federal funds f r o m t h e E n v i r o n m e n t a l Protection Agency G r a n t KO.APO 1443-01. T h e contents d o n o t necessarily reflect t h e views a n d policies o f t h e E n v i r o n m e n t a l Protection Agency, n o r does m e n t i o n o f trade names or commercial products constitute endorsement or recomm e n d a t i o n for use.
Nitric Oxide Reduction with Hydrogen on Mixed Metal Oxide Catalysts George L. Bauerle and Ken Nobe* School of Engineering and Applied Sciences, University of California, Los Angeles, California 90024
Catalytic reduction of NO with H2 on a number of mixed oxide and oxide mixtures has been investigated in a laboratory integral reactor using a feed gas containing 520 ppm of NO and 3000 ppm of H2 in a nitrogen carrier gas. Highest yields of NH3 at 90% conversion of NO or higher were observed with rare earth nickel iron oxide and copper-vanadium oxide, copper-nickel oxides, monel, platinum, palladium, and rhodium. Each of these catalysts showed stoichiometric conversion to NH3 at a specific temperature. Platinum and palladium maintained this stoichiometric conversion up to 500°C; monel, on the other hand, showed a sharp drop in NH3 yield above the temperature of stoichiometric conversion (450°C). Lowest NH3 maxima for NO conversion levels above 90% were observed with ruthenium, mixed ruthenium oxides, and a series of nickel-containing catalysts. When tested in auto exhaust, ruthenium and mixed ruthenium oxides were particularly active and appeared to exhibit true selectivity for reduction of NO to Na.
The production of ammonia during the catalytic reduction of nitric oxide with hydrogen has been the subject of a number of recent papers. Some of the studies have dealt with dry feed gas streams (Shelef and Gandhi, 1972a, b)
while others have dealt with real or simulated automotive exhaust (Jones, et al., 1971; Klimisch and Barnes, 1972; Klimisch and Taylor, 19'73; Taylor and Klimisch, 1973). A few studies have discussed the formation of NH3 using Ind. Eng. Chem., Prod. Res. Develop., Vol. 13, No. 3, 1974
185
