I
D. F. PALAZZO, W. C. SCHREINER, and G. T. SKAPERDAS The M. W. Kellogg Co., Jersey City,
N.J.
Low Temperature Recovery of Hydrogen from Refinery Gases Besides producing hydrogen of any purity for hydrogenation, the plant can be adapted to produce synthesis gas for ammonia
T H E rapid commercialization of catalytic reforming in the past five years has made available large volumes of byproduct hydrogen. According to one estimate, approximately 500,000,000 cubic feet of by-product hydrogen from reformer gases is available daily, and within 10 years this figure will increase to over one billion cubic feet (7). Ammonia from this one source can supply almost twice as much fixed nitrogen as was made from ammonia in 1954-55. Although the ammonia industry is the largest consumer of hydrogen, there are many other outlets, some a t hand and others in process of development. Hydrogen is used now and is expected to be used to a much greater extent in the future for desulfurizing and improving the qualities of feed stocks, cycle oils, and finished products. I t is used in petrochemical operations to hydrogenate unsaturates-for example, cyclohexane can be made from benzene and butanediol from the unsaturated acetylene derivatives. I n the food industry hydrogen is used to remove the acids from fatty oils. Applications that may become more important in the future are processing of shale oil and manufacture of powdered metals. For some processes the hydrogen will be used as it comes from the reformers, but for others partial or complete purification will be required. I n general, the reformer gases are rich in hydrogen, usually ranging from 60 to 95% by volume; the richer gases are characteristic of platinum operations, while the lower concentrations are found in offgases from molybdenum catalyst re-
formers. Impurities are for the most part hydrocarbons-methane through hexane (principally paraffins with some olefins and aromatics and traces of acetylene)-with varying concentrations of carbon dioxide, carbon monoxide, nitrogen, and sulfur compounds. A study of methods of purifying these gases led to the development of a new process, capable of recovering hydrogen of a purity u p to 99.9+%. Essentially it is a combination low temperature condensation and absorption process, in which the bulk of the impurities are condensed out as the gas is cooled. The remaining impurities are then removed in an absorber-stripper system in which liquid propane is used as the absorption medium and a portion of the product as stripping gas. Propane is ideally suited for this application. I t has the lowest freezing point of the paraffinic
Table I. Feed Gas Properties Temp., O F. Pressure, Ib./sq. inch abs. Moisture content
1000
320 Satd. at 320 Ib./sq. inch abs. and 100°
Flow rate, Ib./hr.
F.
7831.5
Composition, Mole yo,Dry Basis Hydrogen 85.2 Methane 8.5 4.4 Ethane Propane 1.8 0.1 Butanes 100.0
hydrocarbons (- 306 ' F.) and a boiling point of -43.8' F., more than 200' F. higher than methane, which is the major impurity that must be absorbed and then stripped from the propane. I t adds practically no impurity to the product gas, unlike a nitrogen wash, which introduces nitrogen in the washing operation. Make-up propane need be added only to compensate for the very small quantity lost in the stripper. Because the propane required for make-up can usually be obtained by fractionating a portion of the hydrocarbons condensed from the gas, the plant may be considered a self-contained unit.
Description of Process The feed used in the design of the plant described is typical of platinum reformer off-gases (Table I). Feed conditions may be varied considerably without making any significant change in the process. Plants treating gases containing as little as 65% of hydrogen by volume have been designed, and lower concentrations could be handled in a similar manner. Although 320 pounds per square inch absolute was selected for the present design, gases available a t lower or higher pressures can be treated in the same type of unit. A few modifications may be necessary to adapt the process for conditions significantly different from those given above, but essentially the process will remain the same. The feed is passed through an alumina drying unit, where the water dew point is reduced to approximately -100' F. (Figure 1). The gas then VOL. 49, NO. 4
APRIL 1957
685
leaving the exchanger it is divided into a product stream and a stripping gas stream. The latter goes directly to stripper E-2,while the product flows through exchangers C-3,(2-2,and C-I in that order. The liquefied hydrocarbon fraction from drum F-4 is flashed to approximately 23 pounds per square inch absolute and returned through exchanger C-4. This fraction is then joined with the stripper tail gas and sent through exchangers C-2 and C-I. The liquid hydrocarbon fractions from drums F-2 and F-3 are similarly flashed to low pressures and evaporated to supply part of the refrigeration in C-1 and C-2. The hydrocarbons separated out in F-I are recovered as a liquid which is returned to exchanger C-1 under pressure. As this stream is assumed to be further processed, a hydrogen product stream, a waste gas stream, and a liquid hydrocarbon stream leave the plant. Table I1 summarizes the temperature, pressure, and composition of these streams.
flows through the first bank of multistream extended surface exchangers and is cooled to -120' F. The point a t which the cut is made may be dictated by specifications on the recovery of a given group of hydrocarbons or simply by a convenient process arrangement, if no requirements are to be satisfied. I n the present general case the cut point was set arbitrarily to obtain a fraction rich in propane-butane. From this fraction sufficient propane can be recovered, if desired, to supply the makeup required in the absorber-stripper. After the first separation in drum F-I, the gas is cooled in progressive steps to -160', -200', and -270' F., while a vapor-liquid separation is made at each of these points. The gas leaving the last separator, F-4,flows to the methane absorber tower, E-I, where the residual hydrocarbon impurities are removed by countercurrent washing with liquid propane. The rich propane (rich oil) is then depressured, heated in exchanger C-3, and stripped of the hydrocarbon impurities by a portion of the product gas in tower
from (2-4 it is important to keep in mind the freezing point of the hydrocarbon stream. Where the condensed hydrocarbon is essentially methane, to keep exchanger from plugging, the feed gas should not be cooled to as low as - 296 O F. -the freezing point of methane-nor within a few degrees of -296' F., as the exchanger surface will be colder than the bulk temperature. A practical lower limit is therefore set for the system by the physical properties of the hydrocarbon impurity. I n the several cases that have been processed, the feed gas temperature entering the absorber has been between -270' and -285' F. Stripping gas rates have varied between 10 and 17% of the product gas. In the plant shown in Figure 1, - 270' F. was selected as the transition temperature between exchanger and absorberstripper system. The stripping gas requirement in this case amounted to about 13.57, of the product gas. For
Table II.
Process and Engineering Features
E-2.
Absorber-Stripper System. By varying the operating conditions of the absorber-stripper system it is possible to remove the hydrocarbons to any degree desired. The decision as to how much hydrocarbon is to be removed in the exchanger system and how much is to be absorbed in the propane can be made only after weighing the effect on the other variables of the system-for example, stripping gas rate, tower sizes?etc. I n selecting these conditions and in setting the exit temperature of the feed
The product gas, containing less than
0.057, methane by volume, leaves the top of the absorber and flows to exchanger C-4,where it is warmed to -259' F. and is then expanded in a turbo expander, J-1. The cold gas, now a t -300' F., exchanges heat with the lean propane, which it cools to operating temperature and is then expanded through a second machine. The discharged gas, now at a temperature of -283' F., is returned to exchanger C-4, where it is heated to -220° F. Upon
Product Streams Product Hydro- Waste gen Gas
Pressure, Ib./sq. inch abs. 320 18 308 Temp., O F. 100 67 67 Flaw rate, lb./hr. 2227.3 4316.4 1287.8 Composition, Mole % 1.0 99.95 48.0 Hydrogen 34.8 4.1 Methane 0.05 14.8 37.3 Ethane 2.4 53.4 Propane Butane __ 4.2 - __
... e . .
. . a
100.0
...
100.0
FEED GAS
c-2 2nd CONDENSER
c* c, c.
37 3 53 4 42 '10-
Ut C
C-4 FINAL CONDENSER
F _ - I. METHANE ABSORBER
380 803
821
12878 -120°F
2320P110 03 MW IOO'F MOL% #/HA 9995 22183 005 90 1 m
2nd EXPANDER
PRODUCT COMPRESSOR
22273
Figure 1.
686
C-3 3rd CONDENSER
INDUSTRIAL AND ENGINEERING
CHEMISTRY
Liquid Hydrocarbon
Low-temperature hydrogen purification plant
E-2 METHANE STRIPPER ~
100.0
HYDROGEN IN THE PETROLEUM INDUSTRY e-4 FINAL CONDENSER
PRODUCT
E - I METHANE ABSORBER
C-6 H, EXCHANGER
i "J
E - 2 METHANE STRIPPER
- 5"
STRIPPER OFF GAS
-
H, RECYCLE COMPRESSOR
LEAN PROPANE .__..
I
I i
- 4*
I
..___ ..__. .___.
.___.
-
C_
"F
._...
i u
I
1I It
-
-
.--_.
STRIPPING
.__..
1st EXPANDER
2 n d EXPANDER
Figure 2.
these conditions the absorber was designed with 10 and the stripper with 24 actual trays. Refrigeration System. The refrigeration is supplied by expanding product hydrogen through an engine. As the amount of hydrogen in the feed decreasm and the concentration of methane and other hydrocarbons increases, it becomes more difficult to satisfy the refrigeration required to operate the plant. This problem may be solved by compressing the feed gas to a higher pressure, providing a larger expansion ratio for the available hydrogen, or by installing a recirculating hydrogen system, which would in effect increase the amount of hydrogen expanded. The decision can be made only after considering the design features and economics. If a higher pressure i s the solution, the plant would look very much iike that shown in Figure 1. Figure 2 shows a typical arrangement of the low temperature section of a plant using an auxiliary recirculating hydrogen circuit. The auxiliary hydrogen is cooled to the temperature of the gas leaving the absorber. I t is then combined with this gas, passed through exchanger C-4, expanded, and used to cool the lean propane. After a second expansion, the gas is split in two streams, one proceeding through exchanger C-G to cool the closed circuit hydrogen, the rest following the same path as the product in Figure 1. This is only one of many arrangements that can be worked out to satisfy the requirements of any
Hydrogen recirculation system
particular case. Although this discussion has considered only hydrogen expansion as the source of refrigeration, the plant may be adapted to other methods. For example, a closed nitrogen cycle may be used, in which the nitrogen is expanded through a machine in a manner similar to that used for hydrogen. Another arrangement might be a cascade refrigeraation system using ammonia, ethylene, and nitrogen as refrigerants. Of the three circuits mentioned-hydrogen, nitrogen, and cascade refrigeration-the first has been found to be most economical. I n a plant similar to that described, although the power requirements were approximately the same, the investment cost for the hydrogen expansion system was the lowest. Nitrogen, Carbon Dioxide, Carbon Monoxide, and Acetylene Impurities. Beside hydrocarbons, nitrogen, carbon dioxide, carbon monoxide, and acetylene are sometimes encountered. Nitrogen and carbon monoxide will pass through with the hydrogen, as only a fraction oE these impurities will be removed in the separators or in the absorber-stripper system. The quantity of nitrogen normally encountered in reformer gases is small, however, and should present no problem in most cases, as it is an inert and will only slightly decrease the partial pressure. If carbon monoxide cannot be tolerated, the gas must be treated by some standard processes, such as absorption in a cuprous salt solution. The treatment for gases containing
carbon dioxide and acetylene will depend on their concentrations. For example if these impurities are present in the gas shown in Table I in concentrations of about 10 p.p.m. or less, the gas may be sent through the cold box directly without danger of depositing solid carbon dioxide or acetylene in the system or accumulating them in the absorberstripper. Because of the small concentrations encountered in regenerative fixed bed or nonregenerative fluid bed reformer gases, these impurities are no problem. Higher concentrations of carbop dioxide will be found, however, in regenerative fluid bed off-gases as a result of entrainment. In that case the gas should be previously treated to reduce the carbon dioxide to 10 p.p.m. Absorption in caustic solution can easily bring the carbon dioxide concentration down to 10 p.p.m. or lower. However, the bulk of the reformer of€-gases are now coming from regenerative fixed bed or nonregenerative platinum units. I n systems such as these there should be no carbon dioxide, carbon monoxide, and nitrogen and perhaps only traces of acetylene-well under the maximum allowable concentration. The removal of such impurities is thus a problem of theoretical rather than practical interest, as far as treatment of reformer gases is concerned. Controls and Instruments. The plant is instrumented to run automatically for the most part. Temperature and pressure indicators and/or controllers are placed at each separator drum and VOL. 49, NO. 4
APRIL 1957
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Table Ill.
Summary of Processing Cost" Cost/Unit
Units Utilities Power Cooling water Steam
600 lb./hr.
Chemicals Propane Start-up H2 Alumina Labor Fixed charges Overhead
a
789 kw. 350 gal./min.
2 . 6 Ib./hr. 40,000 cu. ft./yr. 10,000 lb./yr. 1 man supervision
+
4570 of maintenance of labor
+ 80%
Cents/ 1000 S. Cu. Feet
0 . 7 cents/kwh. 1 . 5 cents/1000 gal. 30 cents/1000 lb.
1.33 0.08
50 cents/lb. 6.8 cents/cu. ft. 32 cents/lb. $2.80/hr. 16%/~r.
0.31 0.08 0.10
0.04
0.75 4.00 1.03 ___
7.72
Based upon 90% stream efficiency.
tower and at points between exchangers to keep track of the operation of the various parts of the unit. The product is analyzed continuously and the lean propane rate is reset, if necessary, to meet the specifications. From time to time, there may be large changes in feed gas composition. These surges will finally be reflected as a temperature change in the last separator drum. T o counteract this a temperature controller, working off the last drum, varies the refrigeration supply to the last exchanger-for example, by changing the expansion ratio or the flow rate of the refrigerant. Pressure drop indicators across the feed passages in the exchangers indicate plugging or build-up of solid deposits. For example, water deposits may accumulate after an extended period of operation in exchanger C-1 of Figure 1. When the ice builds up to a point which prevents normal operation, the exchanger must be derimed. This might be done by injecting alcohol to dissolve the ice. Control valves are jacketed individually to permit separate deriming, if they become frozen or plugged. Local deriming can be carried out by heating a portion of the hydrocarbon waste, or any other suitable stream, in an electrical or steam heater and passing it through the jackets. Start-up and Shutdown. Before
Table IV.
Investment and Processing cost
Processing Plant Capacity, Million S. Cu. Ft./Day 10 5 2 1
688
Plant Cost $810,000 548,000 430,000 400,000
Cost, Cents/ Million S. Cu. Ft.
7.7 10.6 20.6 37.4
the unit can be put into operation, it must be purged with an inert gas to remove the air. The system is evacuated, purged with nitrogen, then evacuated again, and hot. dry raw gas is passed through to dry the system. Finally dry hydrogen is used to purge the unit. Hydrogen is used because of the method used to provide refrigeration in this plant. If feed gas is expanded through the machine, conditions can be obtained under which hydrocarbons are condensed within the expander. To avoid this, bottled hydrogen is expanded and recirculated until the plant has reached operating conditions, propane is introduced into the unit and circulated through the towers, and finally, the feed gas is slowly bled into the hydrogen until the system is operating completely on feed gas. When refrigeration is obtained by nitrogen expansion or cascade refrigeration, the above problem does not arise. After purging, the feed gas is introduced and the refrigeration system is started UP. When the unit is to be shut down, all liquids are first drained from the plant. The propane is sent to storage and the liquefied hydrocarbons are drained from the flash drums, vaporized in heaters, and sent to the fuel lines. The feed gas leaving the cold end of the plant is heated electrically or with steam to 250" to 300" F. The gas can then be introduced a t several points in the circuit and the plant heated to ambient temperature. With the provisions that have been made for local deriming, however, this plant should be able to operate for at least one year without total shutdown. Plant Construction, A plant producing 10,000,000 standard cubic feet per day of product hydrogen can be housed in a cold box approximately 9 feet wide by 15 feet long, consisting of the exchanger section and the tower
INDUSTRIAL AND ENGINEERING CHEMISTRY
section. The first is 11 feet long and 15 feet high, while the tower section is 4 feet long and 27 feet high. The equipment and piping are constructed of aluminum and alloy steels and are housed in a steel casing such as Lindsay structure. The spaces between the casing and the equipment and all the voids between pieces of equipment are packed with an insulating material such as mineral wool.
Economics The approximate investment cost for a plant producing 10,000,000 standard cubic feet per day of 99.9+y0 purity hydrogen would be about $810,000, including the cost of the compressor to deliver the product at the same pressure as the incoming feed gas. The processing cost, which includes utilities, chemicals, labor, operating overhead, and fixed charges, was estimated a t 7.7 cents per thousand cubic feet of product delivered at 320 pounds per square inch absolute (Table 111). I n Table IV are presented the approximate plant investments and processing costs for plants producing 10, 5, 2 , and 1 million cubic feet per day.
Application Besides producing hydrogen of any purity for hydrogenation, the plant described can be adapted to produce synthesis gas for ammonia. A pure nitrogen stream to blend with the hydrogen may be made by catalytically oxidizing the stripper off-gas with sufficient additional hydrocarbons to maintain a hydrogen-nitrogen balance after shifting the carbon monoxide to carbon dioxide and hydrogen. The resulting gas can then be treated for the removal of carbon dioxide and carbon monoxide and blended with the hydrogen to produce synthesis gas. Such a process should be attractive when there is insufficient hydrogen at a site to produce the desired ammonia capacity. From 80 to 90% of the free hydrogen is removed from the gas before it is sent to the hydrocarbon reformers. Thus the size of the reforming section of the plant is decreased and the steam requirement in the shift reactor is minimized, as the hydrogen concentration is greatly reduced.
Acknowledgment The authors wish to thank the M. W. Kellogg Co., New York, N. Y . , for permission to publish the material presented in this paper.
literature Cited (1) Petroleum Processing 10, 1165 (August
1955).
RECEIVED for review October 9 , 1956 ACCEPTEDJanuary 15, 1957
