Synthesis Gas by Partial Oxidation DuBOIS EASTMAN The Texas Co., P.O. Box 300, Montebello, Calif.
M
IXTURES of hydrogen and carbon monoxide are known as synthesis gas. These can be used directly in various catalytic processes t o produce hydrocarbons, methanol, or other synthetic organics. The carbon monoxide present in synthesis gas can be reacted almost quantitatively with steam via the water gas shift reaction t o produce a mixture of hydrogen and carbon dioxide. On the removal of this carbon dioxide, pure hydrogen is available for hydrogenation or ammonia synthesis. Synthesis gas was first produced from coal. A two-stage process was used, first converting the coal t o coke in a coke oven and then reacting the coke with steam in a water gas apparatus to produce synthesis gas. This reaction of coke and steam is highly endothermic. The necessary hcat is supplied by interrupting the gas-making process and blowing the hot coke bed with air. This burns part of the coke and reheats the bed preparatory to the next gas-making cycle. Both the coke oven and the water gas set are operated a t pressures near atmospheric. For this reason the apparatus is large and bulky, and high costs are involved in compressing the synthesis gas to the pressures normally required for use. Another method, which has been widely used for synthesis gas manufacture, is the methane-steam reaction. This reaction is also highly endothermic, but is carried out as a continuous process. This is done by using a catalyst packed in alloy steel tubes which are arranged in a furnace. The equilibrium of the methane-steam reaction requires such high temperatures that, even with the best alloys available, pressures must be held a t relatively low levels. These low pressures again result in bulky equipment and high compression costs. The exothermic reaction of methane and oxygen was studied as early as 1933 by Padovani ( 1 ) who carried out the reaction a t atmospheric pressure in the presence of a nickel catalyst. This same process w-as developed on a commercial scale by Schiller, Bartholome, and Koch a t Opau in Germany. I n the Schiller, Bartholome, and Koch process methane and oxygen were premixed and fed to a reactor through small openings in a refractory burner block, using feed velocities above the rate of flame propagation. As in Padovani’s work, a nickel catalyst was used and the pressure was near atmospheric. Following World War 11, work was undertaken by Hydrocarbon Research on a pressurized version of this process. As in the work of Schiller, Bartholome, and Koch the natural gas and oxygen were premixed and preheated a t velocities above the rate of flame propagation. The reactor itself was internally insulated and packed with a catalyst. The Texas Co. undertook vork on a noncatalytic process using a specially designed burner whirh permitted all the mixing t o be done within the reaction zone. This process was first de-
1118
veloped to use natural gas as a fuel but was later applled to light oil and to heavy bunker-type fuel oils. The first application of this process was a t The Texas Co. Laboratory in Montebello, Calif., where from 1916 through 1954 it was used in a routine manner t o provide synthesis gas for a pilot unit on a modification of the Fischer-Tropsch process. The process was also used a t the synthetic fuel and chemical plant of Carthage Hydrocol. Since that time numerous installations have been made a t synthetic ammonia plants xhere the flexibility of the process with respect to fuel source is particularly attractive. The ability t o produce ammonia from heavy fuel oils by this process has made it possible to manufacture ammonia from easily transported fuel oils a t locations where neither coal nor natural gas is an economically practical raw material. It was immediately apparent that the noncatalytic reaction of methane and oxygen would proceed as a flame. It was also apparent from the general theory of flames that a detailed statement of mechanism was not available. Experimental work soon showed, however, that free oxygen was not a component of the product and that the partial oxidation reaction was characterized by the presence of methane, carbon dioxide, and x a t e r vapor. Although the actual reactions are certainly much more complex, this behavior is consistent with a two-stage reaction mechanism involving complete combustion in the first stage via the highly exothermic reaction
CH,
+ 202
CO:!
+ 2HzO
(1)
Equilibrium is very far to the right of thiq reaction and no measurable oxygen remains. The quantity of oxygen fed is, however, much less than that required t o react with all the methane by Equation 1 so t h a t on completion of this reaction a large surplus of methane remains. This surplus methane can then react with either of the products of the primary reaction as follows: CHa CH,
+ COP 2CO + 2H2 + HQO = CO + 3Hz
(2) (3)
Whereas the primary reaction is highly exothermic, these reactions are both highly endothermic. It is therefore important that the excess heat liberated in the primary reaction be made available by some means of heat transfer t o the secondary endothermic reactions. The balance between these exothermic and endothermic reactions is such that a departure from equilibrium in the secondary stage has the effect of liberating heat. This increases the temperature and tends to accelerate the secondary reactions. TO this extent the reaction system is self regulating
INDUSTRIAL AND ENGINEERING CHEMISTRY
vci. 48,
NO.
7
SYNTHETIC FUELS AND CHEMICALS I n addition to these reactions of methane, the products can react via the feebly exothermic water gas shift reaction
CO
+ HzO = COz f Hn
(4)
and carbon can be formed from carbon monoxide via 2co =
con + c
(5)
Equilibrium calculations of temperature and effluent composition can be based on Equations 2,4, and 5 which include the principal products: CH4, H,, CO, H,O, COS, and C. Values of the equilibrium constants for these reactions have been translated from Wagman ( 3 ) . It has been found that the methane and oxygen requirements of this system can be fairly well approximated by a simultaneous calculation of temperature by heat balance and composition by equilibrium. In these calculations Equation 5 is used only to see that the product mixture is not in the carbon-forming range. Actually the system is rate controlled by the slow secondary reactions, Equations 2 and 3. The temperature is, however, very sensitive to the proportions of methane to oxygen, and the equilibrium is also very sensitive to the temperature. As a result, the simultaneous solution for temperature and equilibrium composition offers a fair approximation of the methane and oxygen demaad of the system. Stated in another way, a relatively large departure from the equilibrium concentration of methane in the product (1% as compared with 0.05%, for example) may still correspond to a good approximation of the methane and oxygen which must be fed to produce a given quantity of hydrogen and carbon monoxide. Since the desired over-all reaction CHa
+ '/no* = CO + 2Hz
proceeds with a doubling in volume, it is apparent that compression costs can be greatly reduced by operating a t a high pressure. If this is done it will be necessary to compress only about half as much gas as would be required if the operation were carried out a t low pressure. Furthermore, the methane is usually available a t a fairly substantial pressure, and the oxygen can be pumped as a liquid before vaporizing. The disadvantages of pressure operation lie in the adverse effect on equilibrium and the severe mechanical problems that are introduced in this high temperature process. Equilibrium calculations showed that a t pressures of about 200 to 1000 pounds/square inch gage the effect of pressure on equilib1 ium could be offset by a moderate increase in temperature and that this increase in temperature could be obtained a t very small cost in increased oxygen. Energy balances are conveniently based on values of the heat content of the constituents expressed as B.t.u./pound mole relative to the elements a t 0" R. In'thermodynamic terms: Fhere
(H? - Hg)
+ AH;
H? = heat content a t temperature, t, ' F.
HI = heat content a t 0" R.
AH,: = heat of formation from elements a t 0'
R.
The 0 superscripts indicate that these values are exact only a t infinite dilution. The errors due to their use a t several hundred pounds pressure are normally unimportant. Values of heat content as defined have been calculated from Rossini ( 2 ) . When dealing with fuel oils, the heat content of the oil, also referred to the elements a t 0' R., is evaluated in the usual way from the heat of combustion. As the ratio of carbon to hydrogen in the fuel is increased a point is reached a t which the quantity of oxygen necessary to react with the carbon is more than enough to bring the product mixture to the necessary temperature. At this point steam can be added to supply part of the total oxygen and to hold the reaction temperature a t the desired level. July 1956
Calculations of heat and material equilibrium such as those discussed will give fair approximations of fuel and oxygen requirements. They may be further refined through the use of reaction rate data which measure the rate of approach to equilibrium in the slow secondary reactions. Experience has shown, however, that more reliable predictions can be made by means of the thoroughly empirical relations discussed below.
Equipment Calculations of the type discussed in the previous section made it clear that equipment must be designed to withstand temperatures above 2000' F. if the methane content of the product were t,o be reduced to a value on the order of 1%. The very severe combination of a pressure of several hundred pounds with a temperature above 2000' F. was met through the use of internally insulated reactors. These are designed with cylindrical carbon-steel pressure shells lined with successive layers of specially selected refractories. I n this process it is possible, through a control failure, to subject the inner wall of the reactor to temperatures approaching 6000" F. However, in normal operations the refractories are subject only to temperatures that are allowable for firebrick. Since the rate-controlling reactions are temperature sensitive, there is a relation between allowable operating temperature and the required residence time. I n practice residence times are selected as required to bring the required temperature within the safe working temperature of the refractory wall. It is also necessary to select values of internal diameter and length such as to provide adequate heat exchange between the exothermic primary reaction and the endothermic secondary reaction. It is essential to avoid contact between hot, unreacted oxygen and the reactor wall, as this is highly destructive. Care must also be exercised to provide for the differential expansion of the refractory and the external pressure shell and a t the same time to prevent the access of hot gases to the metal surface. It is necessary to mix the reactants quickly and thoroughly as soon as they are introduced into the reactor in order to prevent thermal decomposition and coke formation. These reactions proceed very rapidly a t the temperatures used. The mixing process is carried out by means of an alloy burner having carefully controlled dimensions. It would be desirable to use an uncooled refractory burner in order to reduce heat loss and avoid dependence on coolant supply, but no satisfactory refractory has been available. When the product gases are t o be fed to a water gas shift reactor, it is convenient to quench directly with water, so that the sensible heat content of the effluent gas is used to generate the steam needed in the shift reaction. When the hydrogen-carbon monoxide mixture is to be used directly without shifting the carbon monoxide, the heat of the effluent is transferred to steam in a boiler. Care must be exercised in the design of a boiler for this service t o avoid attack through alternate oxidation and reduction. Provisions must also be made to prevent fouling with soot or to remove such soot as may be deposited. After the primary cooling step, any remaining traces of soot are removed by scrubbing. Because of the sensitivity of reaction temperature to the fuel-oxygen ratio, unusual care is required in instrumentation. I t is, in fact, unlikely that the process could be used a t all without its unique arrangement of control instruments. In this service the usual ratio controllers are undesirable since they do not provide a simultaneous record of both flows and both pressures I n the event of an upset condition it is desirable to have this information immediately available so that the operator can determine exactly and quickly where the upset originated. As a result of these considerations, separate pressure and flow controllers are used on each feed stream, and a ratio alarm is in-
I N D U S T R I A L AND E N G I N E E R I N G CHEMISTRY
1119
Natural Gas Operation Table I.
Generator Operation Fuel Natural Oil Gas
Flow rates Natural gas, thous. std. cu. ft./hr. Oil, lb./hr. Steam, lb./hr. Oxygen, thous. std. cu. ft./hr. Dry product gas, thous. std. cu. ft./hr. Operating conditions Pressure, lb./sq. inch gage Oxygen preheat, F. Hydrocarbon preheat, ' F. Product composition, mole % Carbon monoxide Carbon dioxide Hydrogen Nitrogen Methane Hydrogen sulfide Carbonyl sulfide Performance data Oxygen consumption, cu. ft./ 1000 CU. It. Hz CO Cold gas efficiency (heating value of HI CO), % heating value of HC Fed
+
Coal Tar
32.08
...
26.51
428.7 162.7 5.20
4i6:7 218.0 4.57
100.03
21.59
18.73
219 209 827
342 69 712
347 64 750
38.02 2.19 59.54 0.15 0.10
.... ....
47.97 3.65 47.45 0.22 0.26 0.44 0.02
54.36 5.69 38.87 0.82 0.07 0.18 0.01
271.7
252.3
259.0
83.89
86.41
.... ....
+
+
stalled to \Tarn of any deviation from the desired ratio. dutomatic shutoff devices are provided t o stop both flov-s if the ratio departs more than a specified amount from the set value. Platinum, platinum-rhodium thermocouples are used in refractory protecting tubes. Temperature alarms are at,tached to the exterior of the pressure shell. These consist of long, small-bore alloy tubes containing two conductors separated by a temperature-sensitive insulating material. These tubes are stapled spiially to the outside of the pressure shell. A resistance-bridge potentiometer sounds an alarm if any part of the tube exceeds a safe temperature. The methane content of the dry product gas, which is generally the control point in the system, is monitored by infrared analyzers.
Table II. Feed streams Natural gas Methane Ethane Propane Butane Pentane Carbon dioxide Oxygen Nitrogen Total Oxygen Total Input Products Carbon monoxide Carbon dioxide Water vapor Hydrogen Nitrogen Methane Total, Wet Dry
1120
Moles/ Hour
Mole % 81.91 7.01 6.20 1.69 0.71 0.91 0.14 1.43 100.00
69.243 5.926 5.246 1.424 0.596 0.769 0.123 1.209 84.536
38.02 2.19
..
59.54 0.15 0.10
100.00
+
... ...
154.395
106.278
100.231 5.781 30.095 156.930 0.385 _0.266 293.688 263.593
100.231 5.781
...
276.972 35.556 41.968 14.240 7.152
... ... ...
375.888
...
...
...
60.190 313.860
_0.266 106.278
_1.064 375.114
...
+
+
Heat and Material Balances
_69.859 _ _ - _ _..._ _ _ _..._
a .
+
Elemental Balances C H 0 69.243 11.852 15.738 5.696 2.980 0.769
+
+
+
81.86
Typical operating data for natural gas-oxygen, fuel oil-0x3-gen, and coal tar-oxygen systems ale shown in Table I. The results with natural gas were obtained in a generator of commercial size, 11-hereasthe results of fuel oil and coal tar were obtained in a smaller. pilot-size generator. Except for heat losses which are a function of generator size, considerable experience has shown that there are no substantial differences in operation. in generators of different sizes. Material balance and heat balance calculations for the test period on natural gas are given in Table 11. I n this method of calculation the volume of product gas is obtained by carbon balance, and the water vapor content of the reactor effluent iE taken as the average of the value indicated by hydrogen balance and the value indicated by oxygen balance. For the natuial gas operation, the d i y pioduct volume is obtained by dividing the total moles of carbon fed (106.278) by the molal fraction of carbon in the product (0.38025 0.02193 0 00101 = 0 40319). The resulting product volume (263.593) multiplied by the molal composition of the product gives the quantities of all constituents, except water vapor. T n o values for water vapor can be obtained. By subtracting the 1.064 = total atoms oi hydrogen as calculated (313.860 314.924) from the total atoms of hydrogen fed (375.888). a value of 30.482 is obtained by hydrogen balance. Correspondinglg, 11.562 = by subtracting the total a t o m of oxygen (100 231 111.793) from the total atoms of oxygen fed (141.502), a value of 29.709 is obtained by oxygen balance. The average value 29.709)/2 or 30.095 is used. (30.482 The heat balance is obtained by taking the sum of the products of the feed constituents times their respective total heat contents a t the preheat temperature. I n the example given in Table I1 the total heat content of the natural gas stream is given as (69.243) (5.926) (-11 22) (5246) (-9.53), e t c , for a (-1640) total of -1377.7 a t 800" F. The corresponding total a t 900" F. is -1234.9 so that the value for the actual preheat temperatuie of 827" F. is interpolated at -1339.1 thousand B.t.u. per hour. Correspondingly, heat content of the oxygen stieam at 208" F. is given as 325.5 thousand B.t.u. per hour so that the total 325 5 or -1013.6 heat input in the feed streams is -1339.1 thousand B t.u. per hour. In thr samr Jvay, the hrat contrnt
...
... ...
... ...
...
1.538 0.246
...
139.718 141.502
Total Heat Content, Thous. B.t.u./Mole 800' F. 900' F. Total - 16.40 - 14.98 -11.22 -9.53 -8.35 -7.64 -157.38 +9.10 _ _ _ _ +8.86 - 1377.70 200' F. f4.59
...
... - ...
141.888
300' F. f5.31
- 1339.1 at 827' F. =
f 325.5 at 209'F.
- 1013.6
2500' F. ~100.231 11.562 30.095
-8.80 -6.03 -3.79 -2.03 - 156.18 +9.89 i-9.60 -1234.90
-26.30 -134.78 -74.78 +21.27 122.43 ~ +14.63 ~ -2315.3
2600' F. ___ -25.45 -133.36 -73.63 122.05 +23.27 +16.73 ~ ~ _ -2064.0 = -2184.6 at 2552O F.
INDUSTRIAL AND ENGINEERING CHEMISTRY
Vol. 48, No. 7
_
SYNTHETIC FUELS A N D CHEMICALS of the products is calculated to be -2315.3 thousand B.t.u. per hour a t 2500" F. and -2064.0 a t 2600' F., giving an interpolated value of -2184.6 thousand l3.t.u. per hour a t a temperature of 2552" F. The difference between input and output (-1013.6 and -2184.6 or -1171.0) represents the heat loss from the reaction system. Since the heating value of the natural gas fed was 38,444 thousand B.t.u. per hour, the loss is 3.08% and the hot gas efficiency is 96.92y0. The over-all heat balance is therefore:
Heating value of natural gas feed Heatjng value of HBand CO in product Heating value of methane in product Heat loss from reactor (as determined Rhove) Heat aqailable for steam production Total
Thousand B.t.u./ Hour Per Cent 38,444 100.00 81.86 31 ,470 102 0.27 1.171 5;701 38,444
3.08 14.79 100.00
The heating value of hydrogen and carbon monoxide in the cold product, expressed as a percentage of the heating value of the fuel fed (81.86Yo in this case) is identified as "Cold Gas Efficiency" in Table I.
Table 111.
Comparison of Observed and Predicted Oxygen Consumptions
Net Heat Input, Thous. B.t.u./Mole Carbon Fed -13.4 - 13.4 -13.0 13.8 -13.6 - 17.5
-
-17.7 -18.0 -22.5 17.1 -22.4 -22.4 -22.4 -22.4 -22.8
-
Methane Content of Dry Product Gas, Mole 70
Observed Oxygen Consumption, Cu. Ft./1000 Cu. Ft. H2f CO
Predicted Oxygen Consumption (Figure I), Cu. Ft./1000 Cu. Ft. H a + CO
0.07 0.10 0.15 0.28 0.30 0.06 0.10 0.16 0.53 0.98 2.29 2.47 2.51 2.57 2.85
261 253 256 252 252 275 271 269 273 2 53 260 261 265 265 264
260 257 260 251 250 275 271 269 274 253 264 263 263 263 263
There is a relationship between conversion (as measured by the methane content of the product) and operating temperature. Over the normal operating temperature range, the methane content of the dry product gas changes by a factor of two for temperature changes of about 50" F. This relationship is sensitive to residence time, changing about 50' F. for each twofold change in time. The relation of methane content and net heat input to oxygen consumption, shown in Figure 1, is much more general in application. The net heat input is defined as the algebraic sum of the heat contents of the input oxygen and gas and the heat loss, expressed as thousand B.t.u./mole of carbon in the natural gas. This completely empirical correlation has been applied to a very wide variety of operating temperatures, pressures, reactor volumes, and methane contents and predicts values for oxygen consumption within 1% of the experimental values. Typical comparisons are shown in Table 111. July 1956
OXYGEN CONSUMPTION,
Figure 1.
C U F T / U . CU F T
Natural gas operat;on-relation consumption
Hz f CO
of oxygen
The correlation shown in Figure 1 may be expressed by the equation
+
O ~ ( C Uft./1000 . CU. f t . Hz CO) = 196.5 - (17j(log % CHdj
- (3.25) (net heat input)
The process is also sensitive to the ratio of oxygen to natural gas. For an increase of about 5% in this ratio, the temperature increases about 200" F., and the methane content of the product gas declines sixfold. It is this sensitivity of methane content to temperature and feed ratio that has made it desirable to install infrared spectrometers to monitor product streams. Although these instruments could probably be used for control, they have actually been installed as multipoint recorders, following several streams intermittently. Although subject to some control through burner design, there is always some net production of carbon black. This material is produced by the thermal cracking of incompletely mixed gas adjacent to the burner and is largely destroyed by reaction with carbon dioxide and water vapor as the mixture passes through the reactor. It is thermodynamically possible to form free carbon by the back reaction of carbon monoxide to produce carbon and carbon dioxide as the reaction mixture is cooled, but this can be avoided by using a rapid initial cooling rate as the effluent leaves the reaction zone. Residual carbon is effectively scrubbed out of the product gas with water. It is not practical to measure the carbon remaining in the gas after this scrubber, because of the very small quantities and the small particle size. The best evidence of its complete removal is found in the long continued operation of a fixed-bed shift converter on the scrubbed gas. After a year of operation such a converter was found to be free of any appreciable deposits of carbon black. The carbon produced in the gas generation operation is similar in particle size to commercial carbon black, but has a very high oil absorption index.
INDUSTRIAL A N D ENGINEERING CHE.MISTRY
1122
operations in which the oxygen rate is too low, and efficiency suffers because of poor conversion. Substantial quantit,ies of unconverted carbon are produced in this operating range. On the upper limb of the curve, hovever, the conversion of carbon is substantially complete. I n this range a further increase in oxygen rate serves only t,o convert, hydrogen and carbon monoxide to unnanted water vapor and carbon dioxide. Optimum operation with this oil feed then obviously lies nithin 250 to 270 cubic feet of nsygen/1000 cubic feet of Hz CO. From an economic standpoint it is generally desirable to operate in the neighborhood of 250 cubic feet/1000 cubic feet of oxygen consumption and 847, efficiency, although this is subject to some variation n-ith the local costs of oxygen relative to fuel. -4s \vould be expected, the amount of unconverted carbon in an oil generator operation is a function of the oxygen feed rate, the amount of unconverted carbon increasing practically linearly n i t h decreasing oxygen feed rate. Unconverted methane decreases a t a decreasing rate as temperature is increased, the relation being very similar to t'hat described for natural gas. Esperience has sholm that m-ith differences in fuel oil quality and differences in steam ratio, the various relationships betxeen teinperature, unconverted carbon, oxygen consumption, thermal efficiency, and met,hane content of the product gas remain fairly constant,. The last column of Table I shows, for comparative purposes, the results obtained when charging to the reactor a stripped coal tar of very high carbon content. Except for a n even higher rat'io of carbon monoxide to hydrogen, the results are substantially the same as those obtained when charging fuel oil. The oxygen consumption and cold gas efficiency figures are quite comparable. I n most residual fuel oil operations some unconverted carbon is produced, and it is necessary t o install filters to remove this carbon from the circulating water stream. Although this carbon is produced in a particle size comparable to carbon black, it is easily %-etby water, flocculates rapidly, and can be settled completely in about 20 minutes. The settled floc has a loading of about 1 weight 7ocarbon, and it is therefore necessary to filter the stream in most cases. The usual method is t o cool the blowdown stream by exchange wit,h the return n-ater and remove the carbon with a continuous vacuum filter. ?\lention was previously made that most of the sulfur present in the oil feed appears in the product as hydrogen sulfide. Analytical studies show that the distribution of sulfur compounds is very similar to t h a t calculated for equilibrium. One study ehowd, for example:
+
f 30 80
81
82
THEflUAL C f f l C I E N C Y
Figure 2.
83
- HEArINB VALUE
e4
OF U , + C O ,
Fuel oil operation-oxygen
OS
PEfl C E N T O P OIL
and fuel efficiency
Fuel Oil Operation .Llthough similar in many respects to the natural gas operation, the application of this process t o fuel oils is complicated by an additional process variable (steam rate) and by the wide range of fuel oil quality. I n addition sulfur and ash are normal constituents of fuel oils, and their presence needs to be considered. A typical set of data on fuel oil operation is given in Table I. The fuel oil used in this run had the following properties: Gravity, ' API l'iscosity, se:. Furol a t 122" F. Pour point, F. Carbon residue (Conradson), % -Ish content, % Gross heat of combustion, B.t.u. /lb. Ultimate analysis,. 5% .. Carbon Hydrogen Nitrogen Sulfur Oxygen
13.1 405 d0 10.2% 0 06% 18011 55 59 11 38 0 72 1 96 0.35
It is evident from Table I that the general level of performance, as nieasured by oxygen consumption and cold gas efficiency, is very similar t o that for natural gas. The principal points of difference are the higher ratio of carbon monoxide t o hj-drogen for the fuel oil operation and the presence of sulfur in the products. The sulfur compounds present in the oil feed are converted primarily to hydrogen sulfide n-hirh can readily be removed from the product gas by any one of a number of conventional operations, Only very minor amounts of carbonyl sulfide and a trace of carbon disulfide are found in the product gases. Figure 2 s h o w the oxygen consumption (in cubic ieet/1000 cubic feet of HZ CO produced) plotted against cold gas efficiency (heating values of Hz CO, as per cent of heating value of the oil feed) for typical fuel oil operations. Although all these data fall within the range 81 to 847, thermal efficiency, it is evident that a maximum was reached in the neighborhood of 84'3 with about 260 cubic feet/1000 cubic feet of oxygen consumption. The lowcr limb of this curve is representative of
+
1122
+
Distribution of Sulfur, ___Observed 89 11 0 03
H d r o g e n sulfide Carbonyl sulfide Carbon disulfide
yo-
Equilibrium 93
7 0 05
Acknowledgment .Is is commonly the case in n-ork of this nature, many people have been concerned in the pilot plant development of this process. Those most directly insTolved have included L. P. Gaucher, F. E. Guptill, Gordon Kiddoo, J. B. Nalin, C. P. Marion, TV. L. Glater, and D. XI. Strasser.
literature Cited (1) Padovani, C., Francbetti, P., Giorn. C h ~ m .i n d . app2icnta 15, 429 (1933). (2) Rossini, F. D., others, U. S.Bur. Standards, Circ. C-500, Fehruary 1952. and C-461, November 1947. (3) Wagman, D. D., others, U. S. Bur. Standards, Research Paper R P 1634 1.7. Research 34 (February 1945)l. RECEIVED f o r review December 10, 1933.
INDUSTRIAL AND ENGINEERING CHEMISTRY
ACCEPTED March 29, 1056.
Vol. 48, No. 7
