Removal of Argon and Methane from Ammonia Plant Synthesis Gas Carlo G. Alesandrini,l Allen D. Sherman,2 and Scott Lynn* Department of Chemical Engineering, University of California, Berkeley, Calif. 94720
In the temperature range from - 2 0 to 1OO", the solubilities of the components of ammonia synthesis gas in Ar Nz HP, and all increase markedly with temperature. A unique liquid ammonia are in the order CHI type of absorption process has been designed to separate CHI and Ar from ammonia recycle gas. This process would allow the saving of 80-90% of 'the Hz normally lost in the purge gas and greatly facilitate recovery of Ar. The effect of varying design parameters is discussed.
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A m m o n i a is one of the major bulk inorganic chemicals; yearly production in the C.S. exceeds 13 million tons per year. It is synthesized from the elements. The nitrogen is derived from the air and the hydrogen is obtained by reforming methane or other hydrocarbon. Noyes (1967) has described the process, including recent innovations. A flow sheet of the synthesis loop is shown in Figure 1. The feed stream to the synthesis loop typically contains a small amount (1-1.5'%) of methane and argon. These gases are inert in the process, but accumulate with continued recycle. Their presence reduces the conversion in the reactor, necessitating their removal from the system. This is accomplished by purging a fraction of the recycle stream. The resultant loss of hydrogen and nitrogen typically amounts to 5-77' of the feed. X patented process (Bowers, 1959) calls for equilibrating the recycle gas with liquid ammonia a t relatively low temperature (-2°C). The purge is effected by flashing the dissolved gases out of solution a t near-atmospheric pressure. This process depends on having a relatively high (320-70 atm) system pressure and a rather high (25-30'%) level of inerts to dissolve the required amount of inert gases in the ammonia. The loss of hydrogen and nitrogen is still substantial. Twist, et al. (1968), patented a process which calls for equilibrating the recycle gas with ammonia a t a higher temperature (75OC) to take advantage of the solubility behavior to be discussed below. This process also results in only a very limited saving of hydrogen and nitrogen. The solubilities in liquid ammonia of the gases present in ammonia synthesis are found in the literature mostly as binary systems. Alesandrini, et al. (1972), have correlated the data and developed a set of equations which allow prediction of the solubility behavior of the five-component system HrN2-Ar-CH4-?';H3. The system is unusual in that the solubilities of all of these gases increase with increasing temperature in the range from -20 to 100". This behavior forms the basis of the process design discussed in this paper. The evolutionary development of this design has been described elsewhere (Sherman, 1969; Alesandrini, 1971). A patent application on the process has been filed. 1 Present address, University of Trondheim, Trondheim, Norway. 2 Present address, Stauffer Chemical Co., Richmond, Calif.
94804.
Current Practice
Modern large (>800 t / d ) ammonia plants are usually turnkey installations operating under very similar conditions regardless of location. Returning to Figure 1, the system pressure in the recycle loop is usually 200-300 atm. The temperature of reaction is 450-500°C. Since conversion of nitrogen and hydrogen to ammonia is only 20-25% per pass and is limited by thermodynamics, economics dictates that the effluent stream be cooled to condense ammonia and be recycled to the reactor. The cooling takes place in several stages and includes waste heat recovery (through steam generation), heat exchange with the recycle stream, and further temperature reduction by cooling water and refrigeration. These are indicated in Figure 1 with the understanding that more than one stage of any given type of cooling may be involved. The temperatures indicated are approximate. New Process Conception
The simplest version of the new process is shown in Figure 2 as it would fit into the flow sheet of Figure 1. From the converter effluent, 1, a small side stream of vapor 2 is fed to an absorption column. The vapor stream, 3, leaving the top of the column rejoins the recycle stream ahead of the watercooled heat exchanger. The liquid stream, 4, entering the column derives partly from the product ammonia stream, 5, condensed from the recycle gas and partly from liquid ammonia, 7 , which is recycled. The absorption column contains both an absorption section and a stripping section. The liquid stream leaving the column is cooled by heat exchange with the liquid feed and additional refrigeration. This cooling releases the stripping vapor, 6, which is returned to the column. The heat removal section hence serves as a rather unusual type of reboiler. The gas still dissolved in the ammonia is released when the ammonia is expanded to near-atmospheric pressure. It constitutes the purge, 8, from the system and contains all of the argon and methane in the feed plus some hydrogen and nitrogen. A pump returns the recycle ammonia stream, 7, to the system. Process Parameters
The economic usefulness of this process will clearly depend on the sizes of the pieces of equipment required and on the Ind. Eng. Chem. Process Des. Develop., Vol. 12, No. 3, 1973
217
Table 1. Process Parameters for Column in Figure 2
System pressure, 250 atm Nominal column temperature, 90°C Fraction inerts in vapor feed, 0.15 Argon/methane in plant feed, 0.40 Vapor/liquid feed ratio, 0.15 Number of plates (theoretical), 7 Figure 1.
Flow sheet of synthesis loop
L__1
I
I
pwaw
Figure 2. Version of new process
quantities of hydrogen and nitrogen saved. Additional economic incentive will be obtained from the higher argon concentration of the purge if argon recovery is desired. To determine these, a computer program was developed to calculate the behavior of the column as a function of the various input parameters. Plate-by-plate calculations were made using an iterative flash procedure. The correlation of Alesandrini, et ai. (1972), was used to obtain K values for each plate. Convergence for all flows and compositions was to 0.001 (convergence to 0.0001 did not cause significant shifts in calculated behavior). Because of the large number of operating variables and the uncertainty with which equipment costs can be estimated, no attempt was made to optimize rigorously the design of the process modification. Instead, a set of operating parameters has been determined which appears to be well within the realm of economic feasibility, and the manner in which the system responds to changes in a number of the operating parameters has been determined. The set of parameters used for this case is listed in Table 1. The resultant component flows and total flows are listed in Table 11. The flows are expressed in mol/mol of ammonia product and hence are independent of plant size. I n the discussion which follows, the factors influencing the choice of parameters are described as well as the effects of changing them. Feed Gas Composition. The ratio of argon to nitrogen in the feed to the ammonia plant, 0.012 mol Ar/mol Nz, is that of the atmosphere. The ratio of methane to argon in the plant feed will depend upon the operation of the hydrogen plant and will be somewhat variable. The value listed in Table I was chosen arbitrarily. Because methane is more soluble in liquid ammonia than is argon, the ratio of the two in the recycle gas will necessarily be different 2 18 Ind.
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from that in the purge. Because both are inert in the synthesis reaction, it has been assumed that only their sum is important. This was set a t 15%, typical of operation a t 250 atm. Column hessure. The column pressure is set a t the nominal system pressure for the recycle loop. The pressure drop through the column should be less than that through the recycle loop heat exchangers which the vapor flow bypasses, so no vapor compression should be necessary. Increasing the system pressure improves the solubilities of all the gases and reduces the amount of liquid ammonia which must be recycled. Operating a t a lower pressure has the reverse effect. Column Temperature. The feed plate temperature, 90°C, is also the nominal column temperature and is by coincidence the adiabatic saturation temperature of the feed vapor. However, this is not the primary reason for choosing it. The sensible heat of the liquid flowing through the column is much greater than that of the vapor, and the column can be operated a t any temperature in the range 50-1OO0Cby adjusting the heat input a t H-1. Since the solubilities of all the gases increase with temperature, the lowest recycle requirement corresponds to the highest column temperature in the range. Offsetting this advantage is the amount of ammonia required to saturate the vapor. This quantity begins to increase rapidly a t temperatures above 8OoC and represents an increased load on the heat exchangers of the recycle loop. The optimum column temperature is thus a compromise between minimizing ammonia recycle to the column and minimizing heat load in the recycle loop. The column temperature varies longitudinally a few degrees, being somewhat lower on the top and bottom plates. This is the result of changes of ammonia content with position in the column. The heats of solution of the gases in ammonia are not known and were not estimated. Their effect is presumed to be small because the solutions are quite dilute. Column Design. The column consists of seven theoretical stages plus the “reboiler.” The feed enters between the third and fourth stages, counting from the bottom. The absorber stages serve to equilibrate the liquid flow with the feed vapor. If there are too few stages in the absorber, the liquid leaving the feed tray will contain less argon and methane than it could and a higher liquid flow rate will be required. Four theoretical stages were adequate for the purpose. The function of the stripper stages is to strip hydrogen and, to a lesser extent, nitrogen from the ammonia. The stripping vapor is that released in the “reboiler.” The stripping section rapidly reaches a pinched condition and the use of more than three theoretical stages accomplished little additional separation. “Reboiler” Operation. The operation of the “reboiler” is markedly different from that of a conventional absorption column. This difference results from the vapor’s being generated by removing heat instead of adding it.
Table II. Flow Chart for Figure 2 Stream
Converter effluent Vapor feed Vapor return Liquid feed Condenser liquid “Reboiler” vapor Recycle liquid Purge (“3-free basis)
”a
N2
Component flow (n;ol/molNH~ product) HZ Ar
CH4
Total flow
1.142
1.713
5.140
0.741
0.670
9.406
0.035 0.098 1.902 1.064
0.052 0.048 0.003 0.003
0.156 0.155 0.007 0.007
0.022 0.019 0.002 0.002
0.020 0.009 0.005 0.005
0.285 0.329 1.919
0.003
0.034
0.041
0.015
0.022
0.115
0.838
...
...
...
...
0.0072
0.0082
0.0061
The quantity of stripping vapor which can be generated is limited by the finite quantity of dissolved gases in the liquid leaving the last stage of the column. The lower the temperature in the “reboiler,” the larger the flow of stripping vapor will be up to a point. It appears unlikely that a temperature below about -20°C would be practical because of the refrigeration requirement. Furthermore, although solubility data a t temperatures below -20°C are not available, one would especi the solubilities to begin increasing if the temperature were lowered sufficiently. The optimization of the “reboiler” temperature is quite complex if recovery of the argon is contemplated. Reducing the reboiler temperature increases the amount of ammonia recycle which is required. On the other hand, it reduces the loss of hydrogen and nitrogen in the purge. The concentration of argon in the purge is correspondingly increased and its recovery is thereby facilitated. Practical considerations are likely to result in choosing the same temperature for the reboiler as is chosen for the condenser of the recycle loop. Column Feeds. The feed variables which can be set independently are the fraction of inerts in the vapor and the ratio of vapor feed t o liquid feed. Once these are set, total flows are calculated which are those necessary to remove the argon and methane in the plant feed. The fraction of inerts is not independent of the ammonia plant operation, and the choice of this variable will include considerations of the synthesis reaction as well as inerts removal. The higher the fraction of inerts which is allowed, the lower the ammonia recycle required. The vapor/liquid feed ratio chosen also influences the amount of ammonia recycled. The ratio chosen for the sample calculation of Tables I and I1 results in having about 20% of the argon and half of the methane absorbed from the vapor. The argon/methane ratio in the recycle vapor is 1.1, appreciably greater than the ratio of 0.4 in the plant feed and the purge gas. Having the fraction of argon in the column vapor higher increases the amount dissolved and minimizes the ammonia recycle required. Reducing the vapor/liquid feed ratio results in having a larger fraction of the methane absorbed and hence increases the fraction of methane which must be present in the feed. The argon concentration is reduced correspondingly and the amount of ammonia recycled must be increased. However, the lower vapor/liquid feed ratio reduces the quantity of ammonia vaporized to saturate the gas entering the column and hence reduces the heat transfer load in the condensers in the ammonia recycle loop. The
1.081
0.838 0.0152
0.0366
value of the vapor/liquid ratio chosen is great enough that little reduction in ammonia recycle occurs if it is increased. Equipment Design. Column. The column would be operated a t essentially constant flow rates of vapor and liquid, so a sieve tray design would be acceptable. A tray efficiency of 50% was assumed, with a spacing of 0.6 m. The column diameter was sized by the method described by Smith (1963) which considers vapor/liquid ratio, the relative densities, and molecular weights. Because of the very high L/V ratio in the column, the vapor velocity must be quite low. An allowable vapor velocity in the absorber section is in the range 2-3 cm/sec. Because the vapor density is of the order of 0.1 g/cma, the column for a 1500 l / d ammonia plant will have a diameter of the order of 2 m and a height of the order of 10 m for the flows corresponding t o Tables I and 11. Heat Exchange. The heat exchange for this process modification is perhaps best discussed in relative terms. For this purpose it can be compared with the heat exchange requirement of the recycle loop for a conventional plant, such as is shown in Figures 1 and 2. The flow of gas leaving the converter is 9.4 mol/mol product ammonia, with the component flows given in Table 11. It is cooled from the reaction temperature to the refrigeration temperature, a total of about 450”C, and 1 mol of ammonia is condensed. The average heat capacity of this gas is about 7.5 cal/mol “C, and the heat of condensation of ammonia a t 0°C is about 9400 cal/mol. By contrast, the size of the liquid stream leaving the column is about 1.9 mol/mol product. It is cooled about 84OC in its flow through the “reboiler.” The heat capacity of this liquid is about 20 cal/mol “C. Comparison of these heat exchange requirements shows that the heat load in the “reboiler” is only about 8yoof that normally provided in the recycle loop. Since all of the heat exchange to be added is liquid/liquid or liquid/condensing vapor, the additional heat exchange area required will probably be somewhat less than 8% because of the higher heat transfer coefficients. The usual design considerations in which the cost of heat exchange area is balanced against the costs of steam and refrigeration will determine the area finally selected. Pumps. Pumps P-1 and P-2, Figure 2, operate a t system pressure and are required only to provide for the pressure drop through the heat exchangers. P-3 is a turbine or reciprocating piston expander which recovers part of the energy of the liquid leaving the “reboiler” to drive the recycle pump, P-4. In this Ind. Eng. Cham. Process Des. Develop., Vol. 12, No. 3, 1973
219
way, only the modest amount of mechanical energy necessary for P-1 and P-2 is required by this process. Refrigeration Load. The refrigeration load required by H-3 is only a few per cent of that required to chill the recycle stream of Figure 1 to the condenser temperature. Additional refrigeration, not shown in Figure 2, will be required to condense ammonia from the purge stream. Detailed design calculations indicate that the increase in total refrigeration load due to this process modification would be 2-5y0 of the load for a conventional plant. Conclusions
There are two primary benefits to be derived from this process modification-saving of the bulk of the hydrogen and nitrogen presently lost in the purge and beneficiation of the argon content of the purge. For the case calculated, the loss of hydrogen and nitrogen in conventional operation would be 5.0% of the feed. Addition of the process modification would save 91% of the hydrogen and 71% of the nitrogen. (The hydrogen/nitrogen ratio in the plant feed would have to be altered slightly to keep the ratio in the recycle loop a t 3/1.) This saving would be accomplished without changing the operation or capacity of the ammonia plant. The argon content of the conventionally operated plant purge would be 4.9% for the case calculated (ammonia-free basis). The purge from the modified process would be 16.6% h r . The size of the cryogenic plant needed to recover argon from the purge is in approximately inverse ratio to its concentration. The U.S. market for argon in 1970 was 2972 X 10”ft* (3.8 X 106 kg mol) and the rate of growth during the decade 1960-1970 was 18%/year (Chem. Eng. News, September 6, 1971). The argon recoverable from a 1500 t / d ammonia plant is about 0.18 X lo6 kg mol/year, less than 5% of the U.S. market. The cost of argon is very much a function of scale of purchase. At $0.50/100 ft*,the value of the argon recoverable
220 Ind. Eng. Chem. Process Des. Develop., Vol. 12, No. 3, 1973
is of the same order 8s the value of the ammonia which would be saved by this process modification. A detailed cost estimate for this process modification has not been made because, to be meaningful, it would have to be tailored to the requirements of a specific ammonia plant. However, from the relative comparisons made above, one can conclude that operating costs for the ammonia recycle loop would be increased by a negligible amount and that the capital cost should be less than 4% of the total for the ammonia plant and the hydrogen plant combined. The increase of more than 4% in plant capacity, plus the enhanced potential for producing argon, would therefore appear to make this process modification economically attractive. Acknowledgment
The authors are deeply indebted to Donald N. Hanson for writing the computer program for these calculations. Literature Cited
Alesandrini, C. G., PhD thesis in chemical engineering,University of California, Berkeley, Calif., 1971. Alesandrini, C. G., Lynn, Scott, Prausnitz, J. M., Ind. Eng. Chem., Process Des. Develop., 11, 253 (1972). Bowers, F. A. U.S. Patent 2,881,053 (1959). Lynn, Scott, b.S.Patent Application (1970). Noyes, Robert, “Ammonia and Synthesis Gas,” Noyes Develop ment Corp., Park Ridge, N.J., 1967. Sherman, A. D., M S thesis in chemical engineering, University of California, Berkeley Calif., 1969. Smith B. D., “Desi n of hquilibrium Stage Processes,” McGrawHill: New York,%.Y., 1963. Twist, D. R., G. M. Rowell, Brian Hayes, British Patent 1,137,005 (1968). RECEIVED for review September 1, 1972 ACCEPTED December 11, 1972 Sup ort for this project was provided in part by the Patent Fund of t i e University of California. Patent rights have been assigned to the University.
