US 3846278 A
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Nov. 5, 1974 Mi AL P nbbUcT-ni op JET rum.
Filed July 12., 1973 nited' 'States Patent O i@ U.S. CL 208-57 20 Claims ABSTRACT F THE DISCLOSURE Jet fuel is produced from a petroleum fraction boiling from about 135 F. to about 550 F., such as kerosene, by a two-stage hydrogenation process. A platinum group catalyst is utilized in the first stage, a nickel catalyst in the second.
This application is a continuation-in-part of our application, Ser. No. 177,362, led Sept. 2, 1971 now Pat. No. 3,767,562 and copending herewith, the disclosure of which is incorporated by reference herein.
BACKGROUND OF THE INVENTION This invention relates to the production of jet fuel from hydrocarbon feedstocks. In general, a number of methods have been proposed for jet fuel production, from a Wide range of feedstocks. In some processes, various petroleum fractions or products have been subjected to hydrocracking, reforming, alkylation and other processes in various combinations. U.S. Pat. 3,513,085, which discloses jet fuel production from coal liquids and petroleum oils by hydrocracking, solvent extraction, fractionation and hydrogenation is typical of such processes. Other methods of producing jet fuel have involved the hydrogenation of aromatics-containing feeds in various ways, sometimes in combination with such other processes as hydrocracking. For example, U.S. Pat. 3,147,210 discloses the production of jet fuel by catalytic hydrogenation of high boiling aromatic hydrocarbons, preceded by a hydroning or hydrodesulfurization step. The feedstock is desulfurized in cocurrent flow with added hydrogen in the first stage, hydrogen sulfide is stripped after the first stage; the stripped liquid is then subjected to catalytic hydrogenation in countercurrent flow with hydrogen in a second stage.
Detailed specifications for various types of jet fuels have been published by the Armed Forces and ASTM. 'Ihree jet fuels in common use today are those designated LIP-4, JIJ- and ASTM D-1655 Jet A-l fuel. With respect to the more critical properties, the specifications call rfor a maximum sulfur concentration of 2,000 p.p.m. (0.2%) by weight, a minimum IPT Smoke Point of 25 mm. and a maximum aromatics content of volume percent. In addition to methods such as are described in the preceding paragraphs, attempts have been made to use various kerosene fractions directly as jet fuels. However, while these fractions may meet many of the speciiications for such fuels, they often do not meet the IPT Smoke Point specification. Additionally, some kerosenes contain a higher aromatics content than the specifications permit.
Application 177,362 discloses a process for production of jet fuel by a two-stage hydrogenation of a petroleumderived fraction boiling in the range of about 135 F. to about 550 F., prefereably about 300 F. to about 550 F., in which the feed is passed cocurrently with hydrogen over a hydrogenation catalyst in the first stage, and countercurrently to hydrogen over a hydrogenation catalyst, in the second stage. It is disclosed that the hydrogenation catalyst may be any of the Well-known hydrogena- 3,846,278 Patented Nov. 5, 1974 lCe tion catalysts, including such as Raney nickel, or nickel, platinum or palladium, preferably on a support such as alumina, silica, kieselguhr, diatomaceous earth, magnesia, zirconia or other inorganic oxides, yalone or in combination.
The choice of a proper catalyst yfor the conduct of this reaction Will depend, of course, on a balancing 0f the advantages and disadvantages of the catalysts, 4as well as the feedstocks and operating conditions. lPlatinum group catalysts, particularly platinum and palladium, have been found to possess the -advantages of having a comparatively long catalyst life under the operating conditions generally employed, as Well as being less sensitive to sulfur poisoning.
Platinum group catalysts, however, may not always result in as high a conversion rate of aromatics as may be desirable.
It is an object of this invention, therefore, to provide a method for producing jet fuel from a hydrocarbon feedstock without the need for expensive processing steps such as hydrocracking.
It is a further object of this invention to provide a method for producing jet fuel from a hydrocracking feedstock having a boiling range within the temperature range of about F. to about 550 F.
It is a still further object of this invention to provide a process for producing jet fuel from a hydrocarbon fraction boiling substantially within the kerosene boiling range, and more particularly, from a hydrocarbon fraction boiling within the range of from about 300 F. to about 550 F.
A yet further object of this invention is to provide a method for producing a jet fuel with a low aromatics content. Additionally, it is an object of this invention to provide a method for producing a jet fuel which exceeds the minimum IPT Smoke Point of 25 mm.
A still further object of this invention is to provide an optimum catalyst system for such a process.
Other objects and advantages of this invention will become apparent from the specification, drawings and claims thereof.
SUMMARY OF THE INVENTION In brief, the invention contemplates the production of jet fuel by the two-stage hydrogenation of a hydrocarbon feed having a boiling range within the temperature range of from about 135 F. to about 550 comprising the steps of: (a) passing the feed in cocurrent contact with a hydrogen-rich gas through a first hydrogenation zone operated at a temperature of from about 250 F. to about 575 F., at elevated pressure in contact with a catalyst comprising a platinum group metal to at least partially hydrogenate the feed; (b) removing from said first hydrogenation zone a gas phase euent comprising hydrogen and vaporized liquid materials, and a partially hydrogenated liquid hydrocarbon eluent; (c) further hydrogenating the partially hydrogenated liquid hydrocarbon effluent in a second hydrogenation zone operated at a temperature of from about 200 F. to about 500 F. at elevated pressure by passing a hydrogen-rich gas into said second hydrogenation zone countercurrently to said liquid hydrocarbon effluent in contact with a catalyst comprising nickel; and (d) drawing olf from said second hydrogenation zone a gas phase eluent comprising hydrogen and vaporized liquid materials and a liquid phase effluent comprising jet fuel.
The Figure is a diagrammatic illustration of the process of this invention.
DETAILED DESCRIPTION OF THE INVENTION As shown in the Figure, the hydrogenation zones are preferably contained in one hydrogenation vessel, which has the form of a vertical cylinder having dished ends and pressure sustaining walls. The interior of the vessel is divided by horizontal partitions 12, 14 and 24, which are preferably perforated or foraminous plates or the like, into a plurality of chambers or zones including an upper reaction chamber 16, an intermediate vapor-disengaging zone 20, and a lower reaction chamber 18.
The reaction chambers 16 and 18 are packed with hydrogenation catalysts 22 and 23 respectively, as discussed hereinafter. The catalyst 22 in zone 16 is supported on partition 12. The catalyst 23 in zone 18 is supported on a similar partition 24. Partition 24 is preferably spaced somewhat above the bottom of the converter, thus dening the upper boundary of an additional lower charnber or zone 26.
Fresh aromatics-containing feed, such as is hereinafter described, is introduced into the system at line 46, into a hydrogen stream in line 40, and the mixture proceeds in line 40 as indicated by the arrows until it joins line 44, from which may be added a condensed recycle liquid from separator 34. The resulting mixture then passes through line 42 into the top of the hydrogenation vessel, at a temperature of from about 250 F. to about 575 F. and a pressure of from about 400 to about 1500 p.s.i., depending on the boiling range of the feedstock and the severity of the hydrogenation. The lower temperature and pressure correspond to lower boiling feeds and lower severity of treatment.
The mixture of feed, recycle liquid and hydrogen passes downwardly through the catalyst bed 22 in zone 16, under adiabatic reaction conditions in which a substantial amount of the aromatics present in the total liquid charge are hydrogenated to the corresponding naphthenic cornpounds. The reaction mixture which passes out of zone 16 is a two-phase mixture. The liquid phase is a mixture of parains, naphthenes and some unreacted aromatics. The gas phase eiuent is a mixture of hydrogen, inert gaseous impurities, and vaporized liquid hydrocarbons of a composition generally similar to that of the liquid phase euent.
The liquid phase of the eiuent passes downwardly through the vapor-disengaging zone 20 into the second hydrogenation zone 18 (through partition 14, which serves as a distributor plate).
In reaction chamber 18, hydrogen introduced through line 48 passing through chamber 26 contacts the liquid phase etiluent countercurrently, hydrogenating the remaining aromatics to the corresponding naphthenes. The hydrogen is introduced without being preheated, at a relatively low temperature, compared to that of the liquid phase etiluent from zone 16; generally the hydrogen temperature is no higher than about G-120 F.
The liquid portion which emerges from hydrogenation zone 18 is briefly accumulated in chamber 26 of the reactor, permitting disengagement of vapors and sealing the outlet to line 50 to prevent escape of hydrogen. The liquid product is collected in line 50 and contains a very minor portion, generally less than 1.5 volume percent, of residual unhydrogenated aromatics. The gas phase efliuent from hydrogenation zone 18 contains excess hydrogen, inert gaseous impurities, and vaporized hydrocarbons of a composition similar to those contained in the gas phase efiuent from hydrogenation zone 16.
The gas phase effluents from both the rst hydrogenation zone 16 and the second hydrogenation zone 18 collect in vapor-disengaging zone 20. The combined gas phase fraction is Withdrawn through line 28, and is preferably cooled by being passed through heat exchanger or waste heat boiler 52, in which some of the heat is used to produce steam for use in other processing steps, or in other processes, or for general purposes.
The still hot vapor mixture is then passed through line 54, then preferably through condenser 30 in which it is used to preheat the mixture fed to the reactor, then through condenser 32, where the vaporized liquid phase components remaining in the system are recondensed to liquids. The resulting two-phase system, consisting of gaseous hydrogen, inert gases, and reliqueied hydrocarbons, is passed into separator 34, where the liquid and gaseous phases are separated. The liquid phase is passed through line 44 to be mixed with the feed to hydrogenation zone 16 as previously described. The gaseous phase, comprising hydrogen and inert gases, may be partially vented, as through line S6, to prevent build-up of inert impurities in the system.
The remainder, and majority of this gaseous phase is recycled through line 36, to be mixed with the feed to the first hydrogenation zone 16 in line 40. Fresh feed hydrogen gas may be supplied from line 48 through line S8 into the recycle gas, in the event that the recycle hydrogen is insuicient to supply the needs in the rst hydrogenation zone.
An important feature of this invention is a built-in temperature control. Reactions of the type contemplated are exothermic. The production of the desired jet fuel iS favored by low outlet temperatures. Furthermore, runaway reactions much be prevented or coke and/or undesirable side products will be formed. Accordingly, external temperature control means are usually necessitated in processes for hydrogenating aromatics for jet fuel production. The present process, however, provides an inherent temperature control, particularly in the second hydrogenation zone 18. As the hydrogen feed from line 4S passes upwardly through this zone, a portion of the heat present in that chamber is absorbed in the process of sensibly heating the hydrogen.
An additional amount of heat is absorbed by the vaporization of reaction product liquid in zone 18, in an amount suicient to saturate the gas stream emerging from this zone into vapor-disengaging zone 20. Similarly, the temperature in the rst reaction zone 16 is controlled by the absorption of heat in partially vaporizing the liquid feed. The vaporized liquid is removed from the vapor-disengaging zone 20 in conduit 28, as previously described. A similar process for the production of cyclohexane from benzene, with this same built-in temperature control, is described in our U.S. Pat. 3,450,784.
The vaporized hydrocarbons recovered from the vapordisengaging zone 20 and used as recycle comprise partially hydrogenated feed containing up to about 5% aromatics. Because of low concentration of aromatics, the ratio of recycle to fresh feed is less than 1:1, generally in the range of about 0.05:1 to about 0.75: 1, and depends on a number of fractors, including hydrogen partial pressure and purity, desired temperature in the reactor, aromatic content of the feed, etc.
We have discovered that supported nickel-containing catalysts produce higher conversion rates of aromatics and lower residual aromatics content than catalysts of the platinum group. One disadvantage, however, of using nickel catalysts in a process of this type is that such catalysts are quite sensitive to sulfur and tend to become permanently poisoned or deactivated within a relatively short time unless the feed is substantially free of sulfur (less than about 1 p.p.m.). Consequently, to utilize nickel as the catalyst in both zones of the reactor, the feed must be either naturally very low in sulfur content or must be thoroughly desulfurized before utilization. Alternatively, other catalysts may be used, but the conversion levels and residual aromatics content will not be as satisfactory.
We have found, however, that the .advantages of the higher conversion and lower residual aromatics content may still 'be achieved if a supported nickel catalyst is utilized in the second hydrogenation zone 18, with a catalyst selected from the platinum group, more specifically platinum or palladium, with platinum being preferred, being utilized in the rst hydrogenation zone 16. Platinum group catalysts are only reversibly poisoned by sulfur at levels where nickel catalysts would be permanently deactivated, 'and thus, in addition to having a higher sulfur tolerance, also possess a longer catalyst life.
Additionally, the use of the platinum group catalyst 22 in the first, or upper hydrogenation zone 16 permits the processing of feeds containing -appreciably more sulfur than if a nickel catalyst were used in -this zone. With this system feeds may be treated which contain generally up to about 5 p.p.m. sulfur, though sulfur contents as high as l p.p.m. and, in few cases, even 20 p.p.m. may be tolerated, though at these levels catalyst life may begin to decrease. Under ordinary operating conditions it would lbe expected that sulfur leaving the first hydrogenation zone 16 would be in the form of hydrogen sulfide and will be stripped out of the liquid effluent from the first zone by the hydrogen and gaseous products from the second zone 18 before coming into contact with the nickel catalyst 23. However, as an extra precaution catalyst bed 23 may be covered with a layer of zinc oxide which acts as a scavenger of hydrogen sulfide.
The platinum group catalyst is preferably supported on a support such `as alumina, silica, magnesia, zirconia or other inorganic oxides, alone or in combination or on activated charcoal. 'The nickel catalyst may be supported on such materials `as various inorganic oxides, as above, diatomaceous earth or kieselguhr, alone or in combination.
The feed to the process comprises -a petroleum fraction having a lboiling lrange within the temperature range of from about 135 F. to about 550 F. Fractions, Ifor example, with boiling ranges such as 135 F.480 F., 350 F.-510 F. and 300 F.520 F. `are typical of those within this broad range which are suitable feedstocks for this process. The feed can be either a straight run or other petroleum fractions; such fractions as kerosenes, light and heavy naphthas, catalytically cracked cycle oils and furnace oils can be utilized. Particularly suitable is a feedstock generally boiling within the kerosene boiling range, that is, boiling within the range from about 300 F. to about 550 F.
When Isuch a feed is utilized, the vfirst hydrogenation zone 16 is operated at a temperature of from about 300 F. -to about 575 F. and the second zone at 'about 250 F. to about 500 F., within the pressure ranges previously mentioned.
The process of this invention does not accomplish desulfurization forpractical purposes except to the degree mentioned previously, consequently most feedstocks should be at -least partially desu'lfurized prior to being introduced into the process, so that the sulfur content is not greater than about 20 p.p.m., preferably not gre-ater than about 10 p.p.m. and most preferably not greater than about p.p.m. This is generally performed in a separate unit (not shown).
If the feed is desulfurized just prior to its admis-sion into the first hydrogenation zone, it will generally be sufficiently hot that no further heating is required to bring it up to reaction temperature. If, however, the feed has `been obtained from -a simple fractionation process or has been allowed to cool down prior to being passed into this process, or has been in storage, preheating is required. In any case, the hydrogen fed to the first hydrogenation zone 16 must be preheated prior to its introduction into this zone. The liquid recycle to this zone must also be preheated.
The preheating of the hydrogen, and feed if necessary, can be accomplished in a number of ways, and can be performed separately or together. A convenient method, tin this process, is to utilize the heat contained in the vapors in lines 28 .and 54 which have been removed from the vapor-disengaging zone 20. The combined hydrogen (and feed, if necessary) in stream 40, together with recycle liquid from line 44, is passed through heat exchanger 30, in which it is preheated to the desired `inlet tially cooled vapors in line 54. This heat exchange, underv some conditions, may have the additional effect of partially condensing some of the hydrocarbons in the combined vapor stream, facilitating the separation of hydrocarbons for recycle from the hydrogen and other gases, in separator 34.
If the yfresh feed is already sufficiently hot so as not to require prelieating, it should be by-passed yaround the preheater to avoid overheating and undesirable `side reactions. The fresh feed will then enter the system, for example, through line 43 instead of through line 46, or the by-pass can be accomplished in other ways known in the art. In this case, only the hydrogen and recycled liquid hydrocarbons will be preheated.
Alternatively, the preheating of the fresh feed, liquid recycle and hydrogen can be done in separate heat exchangers, and the heated materials mixed before lbeing introduced into the reactor. This separate preheating can be done using any source of -available heat, including the hot vapor mixture in line S4.
The ratio of hydrogen to fresh feed in the mixture fed to reaction zone 16 may vary from a stoichiometric ratio of 1 mole of hydrogen per double bond to as much as about 300% of the stoichiomertic requirement, and the ratio of hydrogen to the liquid material entering re-action zone 18 may vary from -about 0.3 to `about 1.0 moles/mole.
The L.H.S.V. in the first zone 16 is preferably maintained between `about 0.5 Iand `about 6.0, Ibased on fresh fee-d only, while that in the second zone 18 is generally at a higher level. The overall L.H.S.V. is maintained, however, between 0.5 and 6.0.
The temperature conditions in the second zone should be adjusted to maintain the temperature of the liquid product at the outlet between about 300 and about 500 F., depending on the boiling range of the fresh feed, to provide optimum conditions favoring hydrogenation of the aromatics to naphthenes and close equilibrium approach.
It should be noted lthat it is not necessary to saturate =a1l `aromatics in the feed to produce a jet fuel meeting the minimum smoke point requirement. Saturation of of the aromatics is usually more than sufficient to reach this standard; as pointed out hereinabove, the product may have a residual aromatics content of up to 1.5 volume percent. However, much lower aromatics contents can tbe achieved, as illustrated in the Example.
In order to illustrate more fully the n-ature of this invention, `and the manner of practicing the same, the following comparative examples are presented.
A desulfurized straight-run kerosene having a boiling range of 350-500 F. was hydrogenated as shown in the following tabulation, illustrating the beneficial aspects of utilizing a supported nickel catalyst in the second hydrogenation zone.
Top stage Ni/kieselguhr.'
Top stage Bottom starre 60.0- 85.7.
While the above constitutes a description of our invention, it is by no means intended to limit the invention to the specific items disclosed herein, as alternatives and 7 equivalents will readily occur to those skilled in the art. The invention, therefore, is not to be construed as limited, except as set forth below in the claims.
1. A process for producing jet fuels by the two-stage hydrogenation of a hydrocarbon feed having a boiling range within the temperature range of about 135 F. to about 550 F. and substantially free of sulfur-containing impurities, comprising the steps of:
(a) passing the feed in cocurrent contact with a hydrogen-rich gas through a first hydrogenation zone operated at a temperature of from about 250 F. to about 575 F. and at elevated pressure in contact with a catalyst comprising a platinum group metal to at least partially hydrogenate the feed;
(b) removing from the first hydrogenation zone a gas phase effluent comprising hydrogen and vaporized liquid materials, and a partially hydrogenated liquid hydrocarbon effluent;
(c) further hydrogenating the liquid hydrocarbon eiliuent in a second hydrogenation zone operated at a temperature of from about 200 F. to about 500 F. at elevated pressure by passing a hydrogen-rich gas into the second hydrogenation zone counter-currently to the liquid hydrocarbon effluent, in Contact with a catalyst comprising nickel; and
(d) drawing olf from the second hydrogenation zone a gas phase effluent comprising hydrogen and vaporized liquid material and a liquid phase eiuent comprising jet fuel.
2. A process according to Claim 1 wherein the hydrogen-rich gas introduced into the second hydrogenation zone is at a temperature substantially lower than that of the liquid hydrocarbon effluent.
3. A process according to Claim 1 wherein the feed contains up to 5 p.p.m. sulfur.
4. A process according to Claim 1 wherein the feed contains up to p.p.m. sulfur.
5. A process according to Claim 1 wherein the feed to the first hydrogenation zone is subjected to desulfurization prior to being introduced into said zone.
6. A process according to Claim 1 wherein the feed is preheated prior to being introduced into the rst hydrogenation zone.
7. A process according to Claim 6 wherein the gas phase effluents from the first and second hydrogenation zones are combined and passed in indirect heat exchange relationship with the feed to the first hydrogenation zone, thereby cooling said gas phase efliuents and preheating said feed.
8. A process according to Claim 1 wherein the gas phase eluents from the first and second hydrogenation zones are cooled suiciently to condensel the vaporized liquid components thereof, and said vaporized liquid components are separated from the remaining gas components and returned as liquid feed to the rst hydrogenation zone.
9. A process according to Claim 8 wherein a major portion of the remaining gas components is returned to the first hydrogenation zone.
10. A process according to Claim 8 wherein the ratio of recycled liquid to fresh feed is between 0.05 :l and 0.75 :1.
11. A process according to Claim 1 wherein the first hydrogenation zone is operated at a temperature of from about 300 F. to about 575 F.
12. A process according to Claim 1 wherein the second hydrogenation zone is operated at a temperature of from about 250 F. to about 500 F.
13. A process according to Claim 1 wherein the second hydrogenation zone is operated at temperature conditions such that the liquid outlet temperature `from said zone is between about 300 F. and 500 F.
14. A process according to Claim 1 wherein the hydrocarbon feed has a boiling range within the temperature range of from about 300 F. to about 550 F.
15. A process according to Claim 1 wherein the hydrocarbon feed has a boiling range -within the temperature range of from about 350 F. to about 510 F.
16. A process according to Claim 1 wherein the hydrocarbon feed has a boiling range within the temperature range of from about 300 F. to about 520 F.
17. A process according to Claim 1 wherein the hydrocarbon feed is a desulfurized straight-run kerosene.
18. A process according to Claim 1 wherein the catalyst of step (a) comprises a member selected from the group consisting of platinum and palladium.
19. A process according to Claim 16 wherein the catalyst of step (a) comprises platinum.
20, A process according to Claim 16 wherein the catalyst of step (a) comprises palladium.
References Cited UNITED STATES PATENTS 3,450,784 6/ 1969 Reilly et al. 260-667 3,573,198 5/1971 Parker et al. 208-15 3,654,132 4/1972 Christman et al 208-57 3,047,210 9/1964 Hass et al 208-143 HERBERT LEVINE, Primary Examiner U.S. C1. X.R. 208-15, 89