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termined from consideration of equilibrium in the simple C-HpCH4 system. Because of the wide range of variables studied, this work should be useful for design purposes where the question of carbon deposition is important or where gas-phase compositions in equilibrium with solid carbon are needed. The results are applicable to systems comprised of any set of reactants containing only the elements carbon, hydrogen, oxygen, and nitrogen or reactants containing only the elements carbon, hydrogen, and oxygen plus any number of inert components. L i t e r a t u r e Cited Baron, R. E., Porter, J. H., Hammond, 0. G., Jr., "Chemical Equilibria in Car-
bon-Hydrogen-Oxygen Systems", MIT Press, Cambridge, Mass., 1976. Cairns, E. J., Tevebaugh. A. D., J. Chem. Eng. Data, 9 (3), 453 (1964). Mohnot, S. M.. M.S. Thesis, Kansas State University, Manhattan, Kansas, 1977. White, W. B., Johnson, S. M., Dantzig, G. B., J. Chem. Phys., 28, 751 (1958).
Received for review March 16,1977 Accepted January 12,1978 Supplementary Material Available: A thermodynamic analysis, the computational scheme, a tabulation of gas-phase compositions in equilibrium with graphite, and carbon deposition boundaries (47 pages). Ordering information is given on any current masthead page.
Economics of Alternative Distillation Configurations for the Separation of Ternary Mixtures Nickos Doukas and Wllliam L. Luyben" Department of Chemical Engineering, Lehigh University, Bethlehem, Pennsylvania 180 15
The separation of a ternary mixture into three product streams was studied using four different configurations of distillation columns: (1) removing components sequentially in decreasing order of relative volatility in two conventional columns, (2) removing components in increasing order of relative volatility in two conventional columns, (3) a single column with sidestream product, and (4) a prefractionator column followed by a sidestream column. Economic calculations were performed to determine the optimum scheme. Several values of relative volatilities and feed compositions were explored. A single sidestream column was found to be the most economical when relative volatilities are larger than 3:2:1 and when concentrations of the lightest or heaviest components in the feed are less than 10 mol YO.For higher feed concentration, the prefractionator/sidestream column configuration was found to be more economical than the conventional direct or inverted sequences in most cases.
Introduction Optimum configurations of sequences of distillation columns for separating multicomponent mixtures have been discussed in the literature for many years. King (1971) described the use of N - 1conventional (two product) columns to separate N components into N more-or-less pure product streams. Nishimura and Hiraizumi (1971) discussed optimal patterns of sequences of conventional columns. Petlyuk et al. (1965),in a pioneering paper, suggested several alternative configurations for separating ternary mixtures. One-, two-, and three-column schemes were studied for a low relative volatility system (1.2, 1.1,and 1) with symmetrical amounts of the lightest and heaviest components in the feed and over a range of compositions of the intermediate component. Comparisons were made on the basis of the total amount of vaporization required in all columns a t minimum reflux conditions. Capital costs were not considered. Conventional sidestream columns were also not considered. Stupin and Lockhart (1972) discussed one of the schemes proposed by Petlyuk et al. (1965) for a ternary system with high relative volatilities (9,3, and I), a feed that was an equal molal mixture of three components and 90% product purities. No detailed process engineering calculations or economics were presented. Rathore et al. (1974) studied sequences of N - 1 conventional columns with heat integration. Freshwater et al. (1976)
studied essentially the same problem, with four- and fivecomponent mixtures, considering heat integration and conventional columns. Hendry et al. (1972), Rodrigo et al. (1975), Gomez et al. (1976),and Seader et al. (1977) discussed techniques for optimal design of sequences of conventional separation processes for multicomponent mixtures. Tedder (1975), in a study similar to the present one, examined the separation of ternary mixtures of alkanes. No heat integration was considered, but complex designs with sidestreams and multiple feeds were studied. This study is an extension of the work of Petlyuk et al. (1965). Detailed design and economic calculations were performed for four configurations. A ternary mixture of benzene, toluene, and o-xylene (relative volatilities of (6.7,2.4, and 1)) was studied with small amounts of the lightest component (benzene) in the feed and for small amounts of the heaviest component (0-xylene) in the feed. The other two component compositions were assumed equal. The desired product streams were: (1)benzene with 5% toluene impurity; (2) toluene with 5% benzene and 5% o-xylene impurities; (3) o-xylene with 5% toluene impurity. Heat integration was not studied, nor were feeds with small amounts of the intermediate component considered. The pressure in all columns was assumed to be atmospheric since this gave reasonable operating temperatures for both heat input and heat removal. Ideal trays, saturated liquid feeds and
0019-7882/78/1117-0272$01.00/0 0 1978 American Chemical Society
Ind. Eng. Chem. Process Des. Dev., Vol. 17, No. 3, 1978
273
ss
L O F
BENZENE BENZENE PRODUCT ( 5 % Toluene)
11
PRODUCl TOLUENE PRODUCT ( 5 % B4nzene TOLUENE PRODUCT
0-XY LENE PRODUCT
Figure 1. L.O.F. configuration. Figure 3. S.S. configuration with a liquid sidestream draw-off.
H O F
ss
i
BENZENE
V a p o r Sideatream
0-
XYLENE
PRODUCT
TOLUENE
PRODUCT
PRODUCT
Figure 2. H.O.F. configuration. 0-XYLENE PRODUCT
reflux, total condensers, and partial reboilers were assumed.
Figure 4. S.S. configuration with a vapor sidestream draw-off.
Configurations Studied Four configurations of one- and two-column sequences were studied. (1)Two conventional (two-product) columns in which the lightest component is removed as the overhead product in the first column (hereafter called the “Light Out First” configuration-LOF). The bottoms from the first column is fed to the second column which separates the intermediate and heaviest components. See Figure 1.This configuration is probably the most commonly used in industry. Some intuitive reasons why it might be the optimum are the following. (a) Each product stream is taken overhead only once in its appropriate column. Since this means each product stream has to be vaporized only once, the energy consumption would be expected to be lower than with other configurations of conventional columns. (b) Operating pressures commonly are reduced as one moves down the train of columns because the progressively lessvolatile overhead products can be condensed against cooling water at lower pressures. This permits elimination of bottoms pumps and lower temperature heat sources. (2) Two conventional columns in which the heaviest component is removed as the bottoms product of the first column (hereafter called the “Heavy Out First” configuration-HOF). The overhead product from the first column is fed to the
second column which separates the lightest component from the intermediate. See Figure 2. (3) A single sidestream column in which the intermediate component is removed as a sidestream. In this system (hereafter called the “SS” configuration) three product streams are produced from a single column. The lightest component is taken overhead. The heaviest component is removed as bottoms product. The intermediate component is removed as a liquid sidestream from a tray above the feed tray if there is a small amount of the lightest component in the feed. I t is removed as a vapor sidestream from a tray below the feed tray if there is a small amount of the heaviest component in the feed (Luyben, 1966). See Figures 3 and 4. (4) A two-column prefractionator/sidestream column configuration (“PF” configuration). The first column is a prefractionator which makes a rough preliminary split and feeds its overhead and bottoms products to the second column at two different feed tray locations. See Figure 5. The first column overhead contains almost all of the lightest component and about half of the intermediate. This stream is fed to the upper part of the second column. The bottoms of the prefractionator contains almost all of the heaviest component and the rest of the intermediate component. This stream is fed to the lower part of the second column. Thus the function of the
274
Ind. Eng. Chem. Process Des. Dev., Vol. 17, No. 3, 1978 a 1
LOF HOF
ss 0 P F
-
e VI
PRODUCT
E
c
VI 0
V
0
a C E
a
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Figure 5. P.F. configuration. prefractionator is to remove the lightest and the heaviest components from the center section of the second column. The intermediate component is then removed from the second column as a liquid sidestream from a tray located between the two feed trays. The lightest component is the overhead product of this sidestream column. The heaviest component is the bottom product. The products of the second column are exactly the same as in the single sidestream column configuration. In this system, however, the sidestream column has two feed streams, not a single feed as in the previous case.
Computational Techniques The nonlinear algebraic equations describing a steady-state distillation column were solved using a modified Wang-Henke (1966) method. The computer program handled multicomponent systems with nonequimolal overflow, multiple feeds (vapor and liquid), multiple sidestreams (vapor and liquid), and nonideal vapor-liquid equilibrium relationships. The original Wang-Henke method used a successive substitution convergence technique. The “theta” method of Holland and Pendon (1974) was used in this work with much improved results. A typical 30-tray, three-component column could be converged in 10-15 iterations and in less than l/2 s on a CDC 6400 digital computer. Doukas (1976) gives details of equations and numerical techniques employed. Economic Basis Results of any economic evaluation naturally depend on the cost estimates and assumptions made. In this study the cost of each configuration was assumed to be the sum of annual capital costs and operating costs. Annual capital costs were assumed to be 15%of total capital costs of distillation columns and heat exchangers. The graphs of Dryden and Furlow (1966) were scaled up by a factor of 1.64 to determine basic column and heat exchanger costs. Capital costs for piping, insulation, and instrumentation were assumed to add 60% to the basic capital cost. A typical 8-ft diameter column costs $4100 per tray. Typical heat transfer area costs $17.5/ft2. The only operating costs considered in this work were utility costs. All other operating costs, such as labor and raw materials, were assumed to be essentially the same for all configurations and were therefore not included. Estimating the cost of utilities was difficult because the costs of steam a t various pressure levels in a given plant depend on the steam balance and power generation setup in that particular plant. The costs assumed for this study were felt
0
IO
Benzene
20
30
I n F e e d ( m o l e 70)
Figure 6.Total annual cost of the four schemes tested, as a function of the benzene concentration in the feed. to be fairly typical of 1975 prices: cooling water at 90 OF = 6$/10 000 lb; steam at 200 psig = $2/1000 lb; steam at 300 psig = $2.2/1000 lb. Detailed process design information is given in the tables for each case studied. These data should be sufficient to permit reevaluation of the various configurations using different cost numbers. These economic evaluations are dominated by energy costs since utility costs are much larger than capital costs. Each configuration required optimization studies to arrive a t the best combination of reflux ratio, total number of trays, and feed and sidestream tray locations. The conventional two-product columns were optimized using the normal “shortcut” methods. Minimum reflux ratio was calculated from the Underwood equations. Actual reflux ratios were varied over a range of values. Actual number of trays was calculated from the Gilliland correlation. The optimum ratio of actual to minimum reflux ratio was found to be approximately 1.1for all cases. For sidestream columns, optimization was performed empirically by evaluating columns with different numbers of trays, different sidestream and feed locations, and different reflux ratios. Column diameters were determined by assuming a maximum superficial vapor velocity of 2.5 ft/s a t the top of the column. Overall heat transfer coefficients of 80 and 100 B t u h ft2 O F were assumed for all reboilers and condensers, respectively. Inlet and exit cooling water temperatures of 90 and 140 O F were assumed.
Results Two types of feed compositions were studied (1)feeds with varying amounts (5-30%) of the lightest component, benzene; and ( 2 ) feeds with varying amounts of the heaviest component, o-xylene. The concentrations of the other two components in the feed were assumed to be equal to each other. A. Varying Amounts of Lightest Component. Tables I-V and Figure 6 give results for the cases with the concentration of benzene in the feed varying from 5 to 30 mol %. Results show that the single sidestream column is the most economical of the four configurations studied for low concentrations of benzene in the feed (less than 10%).Note that the sidestream is removed as a liquid from a tray above the feed tray. The single sidestream column rapidly becomes uneconomical as benzene concentration in the feed increases above 10%.This is due to the rapid increase in energy costs that result from increases in the vapor boilup requiied to achieve
Ind. Eng. Chem. Process Des. Dev., Vol. 17, No. 3, 1978
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277
Table V. Results for a System with Relative Volatilities 3:2:1 Name of configuration L.O.F. Col. 1
P.F.
S.S.
Col. 2
Col. 2
Col. 1
A. For 5% of the Most Volatile Component in the Feed QD X
Btu/h
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5.7 318.3 5.7 1072 37
11.5 561 12.0 2 041 22 11
11 -
-
29.2 1.85 73 99.5 315 346 268 294 31 060 34 815 120 903 281 094 467 872
15.0 857 15.5 2 716 46 24 29 80.2 120.8 346 268 78 256 374 261 452 517
8.01 401.6 8.13 1905 21
7 -
2.12 83.4 328.5 289.5 26 949 172 074 551 292
11.65 653.2 11.8
2 059 48 39/25 29 60.9 105.4 346 268 68 520 283 749
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Btu/h
AD, ft2 QB X Btu/h AR, ft2 NT NF NS RR Diameter, in. TB,OF T D ,OF Fixed charges on equipment, $ Steam and cooling water, $ Annual cost of system, $
P.F. Col. 2
B. For 20% of the Most Volatile Component in the Feed 12.44 9.60 697.5 469.8 12.60 9.77 2 376.9 1709.9 36 21 17 11
product specifications. The pinch concentration of the benzene in the rectifying section above the feed plate must be kept low enough to keep the concentration of benzene of the liquid sidestream at specification. Therefore, the column vapor rate must be raised approximately in direct proportion to the benzene concentration in the feed. For feed concentrations greater than 10% benzene, the prefractionator/sidestream column configuration is the most economical. Somewhat surprisingly, this configuration is the best even for high concentrations of benzene in the feed (up to 30%). These results imply that many ternary separations may be more economically performed in a prefractionator/ sidestream column setup than in a conventional two-column configuration. The differences in costs are primarily in energy consumption. For example, for the case of 2096 benzene in the feed, the total heat input in the two reboilers in a conventional LOF configuration is 18.8 X 106 Btu/h (Table I). For the prefractionator/sidestream column configuration (Table IV), the total heat input is 13.8 X 1 0 6 B t u h . The differences in capital costs are small. These lower energy costs result from the increased reversibility of the prefractionator configuration. Pinch conditions are avoided in the columns, and mismatches between feed compositions and feed tray compositions are reduced. T o explore the effects of relative volatility on these results, a system with lower relative volatilities of 3, 2, and 1 was studied. For a 5% concentration of the most volatile component in the feed, Table VA shows that the sidestream column is still the most economical, but the cost differential between
8.2 1.9 108.0 110.0 315.6 346.0 268.4 294.4 53 029 36 445 266 783 235 548 591 805
Col. 1 8.81
464.7 9.07 2 173.7 23 10
Col. 2
9.89 554.4 10.00 1751.1 62 5/21 11
1.6 6.3 88.8 97.15 329.6 346.0 280.8 268.4 27 689 74 644 192 094 241 330 535 751
it and a conventional LOF configuration is much less. For a 2W0 concentration of the most volatile component in the feed, Table VB shows that the prefractionator/sidestream column is now the most economical. Therefore, it appears that a single sidestream column is economical when relative volatilities are reasonably high and when only a small amount of the lightest component is present in the feed. For higher feed concentrations of the lightest component the prefractionator/sidestream column becomes the most economical. B. Varying Amounts of Heaviest Component. Tables VI-X and Figure 7 give results for cases with varying amounts of o-xylene in the feed. The single sidestream column in this system has the sidestream withdrawn as a vapor from a tray below the feed tray in order to take advantage of the smaller concentration of 0-xylene in the vapor than in the liquid. For large relative volatilities (6.7,2.4, and 1)results again show that the single sidestream column is the most economical for low concentrations of the heaviest component in the feed (less than 10%). The prefractionator/sidestream column configuration is best for higher concentrations of 0-xylene. Table XA gives results for a system with lower relative volatilities (3, 2, and 1) for 1Wo concentration of heaviest component in the feed. The effect of relative volatilities is very pronounced. Costs are significantly higher, as expected, but now the conventional two-column LOF configuration is the most economical. The cost of the LOF configuration increased from $394 OOO to $654 000. The single sidestream column costs jumped from $334 000 to $1 023 000. The prefractionator configuration costs increased from $364 000 to $711 OOO. Table XB gives results for the same system and for 20% concentra-
278
Ind. Eng. Chem. Process Des. Dev., Vol. 17, No. 3, 1978
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Table X. Results for a System with Relative Volatilities 3:2:1 Name of configuration L.O.F.
P.F.
Col. 2 S.S. Col. 1 A. For 10% of the Heaviest Component in the Feed 18.34 5.92 35.78 1028.41 289.46 2 005.73 18.45 5.96 35.94 2 821.68 1044.38 6 291.35 2 788.13 37 23 45 19 15 39 8 4.7146 0.5621 10.145 130.48 133.55 184.77 300.06 345.96 345.95 268.38 294.34 268.40 72 775 46 121 156 183
Q D X 10-6, B t d h AD, ft2 QB X Btu/h AR, ft2 A s , ft2 NT NF NS RR Diameter, in. TB, OF TD,OF Fixed charges on equipment, $ Steam and cooling water, $ Annual cost of system, $
390 892
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867 271
653 706
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Col. 2
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1 023 454
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711 191
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P.F. Col. 2
B. For 20% of the Heaviest Component in the Feed 16.63 7.19 932.2 351.7 16.78 7.29 2 812.0 1275.7 37 22 19 13
QD X Btu/h AD, ft2 Q B X IO+, B t d h AR, ft2 NT NF NS
Col. 1 10.21 547.1 10.36 2 154.6 23 13
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RR Diameter, in. T B , OF TD,OF Fixed charges on equipment, $ Steam and cooling water, $ Annual cost of system, $
4.80 1.17 124.51 127.15 307.20 345.96 268.37 294.37 66 313 42 846 355 455 175 757 640 361
1.035 3.66 95.82 112.74 321.68 345.93 276.60 268.41 30 735 101 068 219 437 323 807 675 044
F u t u r e Work These studies have investigated only a small area of an enormous field. Many alternative configurations and systems remain to be explored. The system with small amounts of the intermediate in the feed is important and interesting for future study. Also, the performance of a fifth configuration consisting of a sidestream column and a sidestream stripper should be studied. Finally, it would be interesting to determine the effect of changing product purities on all configurations. The dynamics and controllability of these configurations, some of which present potential interaction problems, must be investigated and factored into the selection of the best configuration.
oLOF
Iss f 300'
0
20
10
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in
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F e e d (mole % )
F i g u r e 7. Total annual cost of
the four schemes tested, as a function of the o-xylene concentration in the feed.
tion of the heaviest component in the feed. The conventional two-column LOF configuration remains the most economical.
L i t e r a t u r e Cited Doukas, N., M.S. Thesis, Lehigh University,June 1976. Dryden, C., Furlow, R., "Chemical Engineering Costs", The Ohio State University,
1966. Freshwater, D. C., Ziogou, E., Chem. Eng. J., 11, 215-222 (1976). Gomez, A., Seader, J. D., AIChEJ., 22, (6),970 (1976). Hendry, J. E.,Hughes, R. R., Chem. Eng. Prog., 68 (6),71 (1972). Holland, C. D., Pendon, G. P., Hydrocarbon Process., 148 (1974). King, C. J.. "Separation Processes", McGraw-Hill, New York, N.Y.. 1971. Luyben, W. L., /SA J., 37 (1966). Nishimura, H.,Hiraizumi, Y., Int. Chern. Eng. 11 (l),188 (1971). Petlyuk, F. V., Platonov, V. M., Siavinski, D. M., Int. Chem. Eng. 5 (3),555
(1965).
'
Ind. Eng. Chem. Process Des. Dev., Vol. 17, No. 3, 1978 Rathore, R. N. S., Van Wormer, K. A., Powers, G. J., AlChE J., 20 (3), 491 (1974a). Rathore, R. N. S.,Van Wormer, K . A., Powers, G. J., AlChE J., 20 (5), 940
281
Tedder, D. W., PhD. Thesis, The University of Wisconsin-Madison, 1975. Wang, J. C., Henke, G. E., Hydrocarbon Process., 45 (e), 155 (1966).
Received for review April 29, 1977 Accepted January 26,1978 Presented at 4th International Congress in Scandinavia, Copenhagen, Denmark, Apr 18-20,1977.
(1974b). Rodrigo, F. R., Seader, J. D., AlChEJ., 21 (5), 885 (1975). Seader, J. D., Westerberg, A. W., AlChEJ., 23 (6), 951 (1977). Stupin, W. J., Lockhart. F. J., Chem. Eng. Prog., (Oct 1972).
Kinetics of Thermal Liquefaction of Belle Ayr Subbituminous Coal Donald C. Cronauer,, Yatish T. Shah,' and Raffaele G. Ruberto Gulf Research & Development Company, Pittsburgh, Pennsylvania 15230
This paper presents the results of a kinetic study of thermal liquefaction of Belle Ayr subbituminous coal. The experimental work was carried out in a laboratory-scale, continuous stirred tank reactor. The results for the conversion of coal and the production of pre-asphaltenes, asphaltenes, oils, and gases such as C1-C6, NH3, H2S, CO, COP,and water are given as functions of slurry space time and temperature. A temperature range of 400 to 470 OC and a space time range of approximately 5 to 55 min were examined. All the experimental data were taken at a total unit pressure of 2000 psig and coal-to-solvent ratio of 1: 1.5. Two solvents, hydrogenated anthracene oil and hydrogenated phenanthrene, were investigated. The experimental results were correlated by a kinetic model which assumes the reaction mechanism
/\gases
coal
f
oils
pre-asphaltenes
/
asphaltenes
It is shown that this reaction mechanism (with all reaction rates assumed to be pseudo first order with respect to reacting species) correlates data reasonably well at the temperature levels of 400, 425, and 450 O C for all space times, and at 460 and 470 OC for small space times.
Introduction Recent demands for fuel have increased the importance of coal liquefaction processes. Various processes for coal liquefaction are currently being examined a t the pilot plant level including catalytic [Synthoil, Akhtar et al. (1974a, 1974b); H-Coal, Johnson e t al. (1973); Gulf's CCL] and noncatalytic [Exxon, Furlong et al. (1976) and SRC, Anderson and Wright (1975)l. Very few kinetics data on these processes have been published. A good review of the reported studies on catalytic coal liquefaction has been published by Oblad (1976). The kinetics of liquefaction processes have been examined by several investigators, the most recent studies being those of Guinn e t al. (1975, 1976), Whitehurst et al. (1976), Plett e t al. (1975), Shah et al. (1978), and Reuther (1977). These and similar studies point out a strong need for good kinetic data, which are lacking in most publications. A large amount of published data is obtained either in batch reactors or in pilot-scale demonstration units. These data are, of course, not easily amenable to evaluation of the kinetic rate expressions for the coal liquefaction process. In the present study, we outline a kinetic model for thermal liquefaction of Belle Ayr subbituminous coal carried out in a carefully designed laboratory-scale continuous stirred tank reactor. The effect of space time and temperature on conversion of coal and production of pre-asphaltenes, asphaltenes, and oils is illustrated. A kinetic model which is somewhat more sophisticated than those published in the literature is used to correlate the experimental data. Department of Chemical and Petroleum Engineering,University of Pittsburgh, Pittsburgh, Pa. 15261.
0019-7882/78/1117-0281$01.00/0
Experimental Section Materials. Subbituminous coal from the Belle Ayr Mine, WY, of Amax Coal Co. was used. The proximate and ultimate analyses of this coal are given in Table I. I t is noted that the ultimate analysis was performed on samples of pulverized coal with subsequent correction for moisture content. Table I. Analysis of Belle Ayr Coal Samples As received Proximate analysis,a Wt % 30.25 0/6 Ash 7.01 % Volatile 28.26 % Fixed carbon 34.48 100.00 Heat of combustion. Btullb 8141 % Moisture
Dry basis
... 10.05
40.52 49.43 100.00 11 671
Ultimate Analysis,b Wt % on a Dry Basis Carbon 69.3 Hydrogen 4.3 Nitrogen 1.0 Oxygen, determined 19.9 Difference 14.5 Sulfur, organic 0.45 Pyritic 0.06 Sulfate 0.00 Total 0.5 Ash, total metals 10.3 5.0
Total' 100.0 Average of two spot samples. * Average of samples taken during experimentation. Sum of C, H, N,ODiff, ST, ash = 100.0.
0 1978 American Chemical Society
