US3328288A - Production of supersonic jet fuels - Google Patents

Description

FUEL. GAS

June 2?, E967 C. w. sTREx-:D 323,233

PRODUCTION OF SUPERSONIC JET FUELS Filed Sept. 26, 1963 Iv a J5 HYDROTREATER |4 3 u Nouavulxs T souvwouv z g o L@ HBXW INVENTOR. m CARL w. STREED m r (D N /ww 2K ATTORNEY United States Patent O 3,328,288 PRODUCTION F SUPERSONIC JET FUELS Carl W. Str-eed, Haddoneld, NJ., assigner to Mobil Oil Corporation, a corporation of New York Filed Sept. 26, 1963, Ser. No. 311,833 12 Claims. (Cl. 208-89) The present invention relates to a process for the manufacture of hydrocarbon fuels for jet engines. It is particularly intended for the production of fuels for jet engines of aircraft operating in supersonic speed ranges. However, the fuels produced by the process described herein may also be used in operating turboprop and turbojet airplanes, at subsonic speeds as well as stationary turbojet engines. Certain combination claims herein include procedures which are also described and claimed in my concurrently filed application Ser. No. 311,869.

Although turbojet and turboprop engines may be operated with a variety of hydrocarbon fuels, specially refined kerosines are preferred for the purpose and suitable specifications for aviation kerosine jet fuels JP-4 and JP-S are set forth in United States military specifications MIL-J-5624E.

In the case of supersonic jet planes operating under more severe conditions than subsonic jet aircraft, especially in regard to extremes of temperature, more severe specifications are desirable for optimum performance than those established for the slower planes. In general, fuels boiling Within the range between 300 and 600 F., and preferably between 350 and 550, of high parainic contents (including isoparans) are desired; and they should not contain more than small amounts of aromatic hydrocarbons (for example, 5% or less). In view of the intense cold at the high operating altitudes of Such aircraft, such fuels should have freezing points below about -20 F., and preferably below about -30. The supersonic fuels desirably comprise relatively high proportions of isoparains which have lower freezing points than paraffins containing the same number of carbon atoms per molecule; and they should not have substantial amounts of benzene, o-xylene and p-xylene for these compounds have relatively high solidication points. The net heat of combustion should be at least 18,890, British thermal units per pound (B.t.u./lb.), and a high 1uminometer number (ASTM test D 1740-60T) of at least about 90 vis also sought to avoid excessive smoking during operation. ln addition, the fuel should have a high flash point of at least 150 F. in combination with a low vapor pressure of less than 50 pounds per square inch (p.s.i.) at 500 F. High grade fuels for supersonic aircraft and their manufacture are described in detail in the copending application Ser. No. 173,044, filed Feb. 13, 1962 of Halik, Smith and Streed, which application was later abandoned.

In the manufacture of jet fuels from kerosines, the principal conversion of hydrocarbons involves the dehydrogenation of naphthenes to aromatic hydrocarbons with libertaion of hydrogen, and secondarily isomerization reactions. The aforesaid conversion may be carried out at elevated temperatures and pressures in the presence of large amounts of hydrogen in the form of hydogen-rich refinery gases over a catalyst of the same type employed for reforming naphthas to upgrade their octane ratings.

3,328,288 Patented June 27, 1967 ice Although the preferred kerosine feedstocks have a high content of normal and isoparaliins, diiiiculties are often encountered in converting them in the manner described as a result of an inadequate or borderline production of hydrogen during the catalytic conversion. This particular problem is of no great moment in refineries that have two or more naphtha reformers as one of the reforming units can easily generate enough by-product hydrogen in the routine reforming of naphtha to make up any hydrogen deficiency encountered in another reformer used for jet fuel production. However, Very few refineries have more than a single catalytic reforming unit, and commercial hydrogen from other sources is relatively expensive. In addition, it is usually necessary to pretreat the kerosine feed with hydrogen in order to avoid poisoning of the dehydrogenation catalyst by compounds containing sulfur, nitrogen or metals. Such a hydrogen treatment consumes some hydrogen and this must be obtained from the dehydrogenation reaction if the overall process is to operate in a self-contained manner.

In the conversion of kerosine to jet fuels by such processes, it has been found that a production of about 300 standard cubic feet measured at 60 F. and standard atmospheric pressure (scf.) of hydrogen per barrel (42 U.S. gallons) of normally liquid feed charged to the conversion zone constitutes a border line operation and that producing about 340 s.c.f./b. or more is necessary to avoid hydrogen deficiencies. Approximately, onehalf of this hydrogen is needed to make up for leakage, mechanical losses and the hydrogen dissolved in the liquid product, while the balance is required for any preliminary purification of the feed with hydrogen that may be necessary and for venting from they system to avoid accumulating excessive amounts of inert gases in the recycle gas.

An object of the invention is to provide an improved process for the manufacture of jet fuels.

Another object of the invention is to provide adequate production of hydrogen in the manufacture of supersonic jet fuels from highly parafnic feedstocks.

A further object of the invention is to improve yields in the manufacture of jet fuels.

Still another object of the invention is to provide a self-suicient conversion process for the manufacture of supersonic jet fuels from kerosines of a highly parafinic nature.

A still further object of the invention is to increase the productive 'capacity of catalysts and conversion vessels in the manufacture of supersonic jet fuels.

Other objects and advantages of the invention will be apparent to those skilled in the art upon consideration of the detailed disclosure hereinafter.

Unless otherwise indicated herein, all temperatures are expressed as degrees Fahrenheit F.), pressures as pounds per square inch gage (p.s.i.g.), boiling points or ranges in degrees Fahrenheit at atmospheric pressure by the ASTM procedure, proportions in terms of weight and space velocities as volumes of normally liquid charge per total volume of catalyst bed or beds per hour (l.h.s.v.). Further the expressions major and minor are used to describe more than half and less than half respectively of any given quantity.

The present invention is a process for the production of jet fuels which includes the steps of removing at least a substantial proportion (preferably a major proportion) of the aromatic hydrocarbon content of a hydrocarbon feedstock (preferably kerosine) containing paraftinic, naphthenic and aromatic hydrocarbons, thereafter subjecting the reduced stock (that is the dearomatized portion or remainder of the feedstock) to dehydrogenation conditions in the presence of a recycled hydrogen-rich gas anda dehydrogenation catalyst for naphthenes, separating the normally gaseous and the normally liquid fractions of the `dehydrogenation reaction ellluent, recycling a substantial portion of said gaseous fraction to the dehydrogeneration reaction as the hydrogen-rich gas, removing at least a substantial portion of the aromatic hydrocarbons in said normally liquid fraction and withdrawing the remainder of said liquid fraction as a product of the process.

In particular, the present invention is directed to the use of the aromatics removal step for high boiling kerosine (e.g. 300 to 600 F.) fractions prior to subjecting such kerosines to dehydrogenation, as described in copending application Ser. No. 173,044 under conditions of temperature, space velocity and pressure correlated to effect the desired naphthene dehydrogenation with substantial conservation of the paraffinic hydrocarbon content of the kerosine feed to the dehydrogenation reaction. The instant process is especially suited for the production of the supersonic jet fuels. Excellent results are obtained with kerosines which have low amounts of naphthenes and high contents of paraflinic hydrocarbons, such as those derived from the Qatar and West Texas Bright crudes.

Other more specific embodiments of the invention relate to blending a kerosine of relatively low naphthene content with a hydrocarbon distillate of substantially higher naphthene content to further improve the Production of hydrogen in the aforesaid process; and fractonating said liquid product fraction prior to the removal of the aromatic hydrocarbons produced in the dehydrogenation reaction from that product fraction in the event that said blending stock does not have a boiling range suitable for jet fuels.

Still other features include the use of kerosine feedstocks containing less than about naphthenes by volume. The generation of hydrogen at a rate of at least 340 s.c.f./b. in the conversion step, the use of noble metal catalysts and especially those containing between about 0.1 and 1.0% or more of platinum supported on a particle form alumina carrier, hydrotreating the charge stock with hydrogen-rich gas prior to the principal conversion reaction as well as conducting the principal conon extraneous sources of hydrogen, a number of other advantages result from practicing the process of the present invention. A more highly parafinic feedstock may be employed than was heretofore possible in t-he absence of outside sources of hydrogen and this naturally produces a higher yield of paranic jet fuel. Less severe dehydrogenation-isomerization reaction conditions may be employed because of the reduced proportion of aromatic hydrocarbons in the charge; consequently, there is less cracking and loss of paraffinic hydrocarbons in the kerosine boiling range. Also, the productive capacity of the catalysts and processing equipment is very considerably increased by the prior elimination from the feedstock of most of the aromatic compounds which are not convertible into generally desirable components of jet fuels. Finally, an improvement in luminometer number of the jet fuel has been observed.

Any of the known methods of removing aromatic hydrocarbons from mixtures of these compounds with naphthenic and paraffinic hydrocarbons is suitable for the present purposes. For example, the hydrocarbon mixtures may be subjected to extraction with vari-ous solvents, azeotropic or extractive distillation, treatment with strong sulfuric acid or oleum or adsorption on silica gel. The same or different removal procedures may be employed for the initial and the final steps of removing aromatic hydrocarbons in the instant process. Inasmuch as the aforesaid techniques are well known in the art, it is not necessary to describe them here other than to mention that extraction with sulfur dioxide is generally recommended for both removal steps. A major proportion of the aromatic compounds present may be removed in each of these steps, and it is preferred that this removal be as complete as is consistent with economic processing. To obtain a superior product, it is desirable to produce a jet fuel containing less than about 5% of laromatic hydrocarbons and preferably less than about 3% by volume.

The reduced kerosine feedstock is subjected to conditions conducive to the dehydrogenation of naphthenes to aromatic hydrocarbons and also the isomerization of parafns to isoparafins over a suitable catalyst for such reactions in the presence of substantial amounts of hydrogen. This may be carried out in a catalytic reformer of the type used for upgrading naphthas by dehydroaromatization and cyclization reacti-ons under somewhat different conditions than are employed here.

The principal feedstock is preferably of the kerosine type which may be described for the present purposes as a petroleum distillate boiling within the range of about 300 and 600 F. and preferably between about 350 and 550 F. As described hereinafter, the feed may be supplemented by liquid hydrocarbon mixtures of similar or different boiling ranges for certain purpose; and material outside the kerosine volatility range is removed subsequently by a suitable fractionation procedure. The present process is intended for principal feeds containing substantial -amounts of aromatic hydrocarbons, as for instance at least about 8% by volume and usually at least about 15% by volume. In addition, even though the instant method is designed for stocks of relatively low naphthenic content, the material eventually charged to the dehydrogenation reaction should contain substantial quantities of naphthenes, for example, at least about 25% and preferably at least about 28%, of the total liquid volume of normally liquid hydrocarbons in the charge to the dehydrogenation step.

A large amount of hydrogen is also charged to the conversion reactor in Ia recycled gaseous mixture. This may range from about 1000 to 15,000 s.c.f./b. of hydrogen per barrel of normally liquid feed to the dehydrogenation zone and about 4000 to 10,000 s.c.f./b. is usually preferred for the purpose. Although the hydrogen is typically in admixture with the substantial amounts of normally gaseous hydrocarbons, especially methane, it will be appreciated that these rates refer to only the hydrogencontent of these gaseous mixtures. The purpose of this high concentration of hydrogen in the conversion zone .is to prolong the life of the catalyst between regenerations, and it has little effect upon the production of hydrogen.

The dehydrogenation-isomerization conversion is conducted in the vapor phase at elevated temperatures and pressures. Accordingly, the mixture of liquid hydrocarbons and hydrogen-rich gas is preheated to a temperature of about 800 to 1000" F. in a conventional type heater or furnace. The overall conversion is of an endothermic nature, and the temperature of the gaseous effluent from the reactor is maintained between about 640 and 980 F., the range of about 780-850 F. being preferred for optimum hydrogen production. If the temperature of the reaction effluent is raised gradually through the 640-980" F. range with other reaction conditions held constant, first the hydrogen production increases, then it reaches a maximum and finally decreases in the upper part -of said range. Temperatures that are too low (insufficient conversion severity) require lowering space velocities so much that the productive capacity of the equipment is unnecessarily reduced; moreover, they promote the unwanted hydrogenation of aromatic to naphthenic hydrocarbons with an attendant high consumpton of hydrogen that is also undesirable. Unduly high temperatures (excessive conversion severity) tend to promote the formation of olefins and the cracking of hydrocarbons. Olefin formation is particularly undesirable here since these olefins are subsequently hydroigenated, thereby consuming extra hydrogen; moreover, the cracking removes valuable hydrocarbons from the desired boiling range.

The partial pressure of hydrogen in the conversion reaction may be as low 4as about 20 or as high as about 1400 p.s.i. .and controlling it within the range of abou-t 100-500 p.s.i. is especially `recommended for the instant process. If the hydrogen pressure is gradually decreased within the range of about 500 to 100 p.s.i. with other reaction conditions constant, the hydrogen make increases fairly rapidly for a while and then at a slower rate; however, the catalyst life -decreases if the hydrogen partial pressure is allowed to drop too low in reducing the total reaction pressure. The total reaction pressure, which is of less significance, may =be within the range of about 30 to 2000 p.s.i.-g.

The over-all space velocity (LHSV) of the gaseous charge through the dehydrogena-tion catalyst bed or beds may range from about 0.5 -to 60, and it -is usually set at a value within the range of about 1-20. An excessive space velocity for any given temperature within the stated range produces an insufiicient conversion severity with attendant lowering of both naphthene dehydrogenation and hydrogen production; whereas an inadequate space velocity results in both an excessive conversion severity and an uneconomically low jet fuel production rate.

While any contact material capable of catalyzing the dehydrogenation of naphthenes to aromatic hydrocarbons may be employed in the present process, it is desirable that it serve the dual purpose of also catalyzing the isomerizat-ion reactions. Paraffins are converted to isoparaffins by adding methyl groups or shifting such radicals along .the chains of carbon atoms, and cyclopentanes are isomerize-d to cyclohexanes which are then dehydrogenated to aromatic hydrocarbons. It is also preferable that the ca-talyst have little or no activity of the type that promotes undesired side reactions to any substantial extent including the `dealkylation of aromatic hydrocarbons and the cracking of Iany hydrocarbons to smaller molecules. Ane of the benefits of this invention is high conservation of parafiinic hydrocarbon content Wi-thout Ireduction of the number of carbon atoms per molecule thereof, so substantial cracking -is particularly undesirable here.

Accordingly, such catalysts as tungsten and/or nickel sulfides on kieselguhr, 'oxides of chromium on alumina, etc. may be used in the instant conversion, but noble metal catalysts of the platinum series including platinum, palladium, rhodium, etc., are greatly preferred; and these are desirably dispersed in the finely divided state on inert carrier materials in particle form such, as the various aluminas (especially the eta, gamma and chi varieties), low activity silica-alumina, etc. The carrier may contain halogen, such as chlorine or uorine, in small quantities that do not exceed the noble metal content and preferably amount to less, as the halogen component provides isomerization activity. Low activity silica-alumina also catalyzes isomerization or alternatively the catalyst may contain a constituent such as sulfided molybdena for the purpose.

The conventional catalysts for reforming naphthas in the presence of hydrogen by dehydr-oaromatization, cyclization, etc. provide good results here. Moreover, these possess the advantage of enhancing the exibility of operations in a refinery because the same catalyst and reactor may be used alternatively either for reforming naphthas or for producing jet fuels. Accordingly, platinum reforming catalysts containing about 0.1 to 1.0% of platinum supported on either eta or gamma type alumina of a particle side within the 1/16 to 3/16" range are especially recommended for the instant process. Although moving bed or liuidized catalysts may be utilized in the present state of the art a fixed catalyst bed is considered more economical for the dehydrogenation-isomerization reactions of this invention. y

Various conversion reaction conditions mentioned hereinbefore are interrelated, and a change in one may require adjustment of another reaction condition. Accordingly, hydrogen is generated at a rate of 340 s.c.f./b. or morea rate higher than is commercially feasible with the raw charge stock-in the conversion of a substantially dearomatized charge stock by controlling the conversion severity in relation to the naphthene content of the charge (which contains at least about 25 volume percent of naphthenes) by selecti-on and adjustment Within the stated limits of certain reaction conditions, especially temperature, space velocity and selection of the catalyst.

The raw petroleum distillates commonly employed as feedstocks usually contain sufiicient amounts of sulfur and/or nitrogen in organic components to temporarily poison or deactivate many dehydrogenation catalysts, including platinum and palladium. In such instances sulfur and/ or nitrogen removal is a necessity for economical processing. This may be accomplished in conventional fashion by preheating the liquid charge along with about 190 to 3000 s.c.f./b. of hydrogen in a recycle Igas to a temperature between about 550 to 850 F. at a pressure within the range from to 1000 p.s.i. and space velocity (LHSV) of 0.5 to 10. The preheated charge in the vapor state is passed through a reactor containing a bed of a typical desulfurization catalyst, as fo-r instance, cobalt molybdate supported on gamma alumina to convert organic sulfur in the charge to hydrogen sulfide and organic nitrogen to ammonia. Metal contaminants in the feed are also removed by deposition on the desulfurization catalyst. After cooling the efliuent, the hydrogen sulfide and/or ammonia are stripped from the purified liquid. The hydrogen consumption in this pretreating or hydrogenation step typically runs between 25 and 150 s.c.f./b. of liquid charge depending on the initial sulfur and nitrogen contents thereof and to some extent, on the nature of the hydrocarbons present. So most of the hydrogen charged to the hydrotreater is conserved by recycling it from the hydrotreater eiuent after purification.

In another modification of the instant process, a kerosine of low naphthene content is blended with a petroleum Adistillate containing a relatively large proportion of naphthenic hydrocarbons, as for example, 32% or more on a volumetric basis. Such blending is especially desirable Where it is desired to manufacture some extra hydrogen for use elsewhere in a refinery or to simultaneously produce a partly reformed naphtha that is an excellent mot-or fuel blending stock. Also, the blending stock may be selected with a View to the selective production of certain aromatic hydrocarbons such as benzene, toluene and xylenes in the conversion reaction from cyclohexane, methyl cyclohexane and polymethyl cyclohexanes respectively in the blending stock.

The initial removal of aromatic hydrocarbons from the feedstock and any blending operation and hydrogen pretreatment (hydrotreating) are all carried out prior to the dehydrogenation-isomerization conversion. However, the order or sequence in which these preliminary process steps are conducted may be varied to meet various objectives, It is generally preferable to perform any blending first, then remove the aromatics and finally effect any necessary hydrotreating in preparing the dehydrogenation charge. However, the hydrogen treatment should desirably precede the -dearomatization step when it iS desired to recover aromatic hydrocarbons of superior purity; whereas the reverse procedure will deliver a smaller and more pure hydrocarbon stream to the hydrogen purification operation and thus reduce the load there. Depending on the quantity of the blending stock, its content of naphthenic and aromatic hydrocarbons and its freedom from dehydrogenation catalyst poisons, it may be advantageous to carry out the blending'after the first removal of aromatics and either before or after any hydrotreating step.

The reaction effluent of the dehydrogenation-isomerization conversion is subjected to a fractionation step, after the hot gaseous efuent has been at least partially cooled, In cases where all of the normally liquid components of the charge to the conversion reaction are within the kerosine boiling range, the gaseous efii-uent is cooled almost to ambient temperature, say within 30 thereof, and then separated into gas and liquid phases in an ordinary separator.

Whenever either a blending stock or the principal feedstock boils at least partially outside of the kerosine range, a somewhat more complex fractionation is necessary. Fractional distillation is recommended for the purpose; and the conversion reaction products are split into three or more fractions. The gas phase rich in hydrogen is taken overhead and a dehydrogenated kerosine cut boiling at about 300-600 F., and preferably between about 350 and 550, is taken off for further processing into a jet fuel. In addition, cuts are taken off `boiling below or above said range, or both, depending upon the composition of the material charged. For example, a heavy naphtha of 400 F. end point may be utilized as the blending stock, and a lighter naphtha (350 E.P.) that has been reformed in part during the conversion treatment may be taken as a by-product from a level higher on the fractionating tower than the kerosine product fraction. In another embodiment, wherein the charge includes heavier hydrocarbons, a gas oil fraction may be withdrawn as the towers bottoms. If the volatility range of the total liquid charge is very broad, both naphtha and gas oil by-product streams may be taken off simultaneously. When such fractional distillation is employed, 'the effluent gases from the converter need not lbe cooled to the atmospheric temperature level.

For a better understanding of the nature and objects of the present invention, reference should be had to the detailed disclosure hereinafter taken in conjunction with the accompanying drawing which is a flow sheet or schematic representation of a system suitable for the practice of the present process. For simplicity many conventional elements including valves, regulators, instruments, etc. have been omitted from the drawing.

Turning now to the flow sheet to illustrate one embodiment of the process, a kerosine feedstock of the type described herein is charged by means of conduit 2 to the mixing tank 4. Through the valved line 6, a naphtha with a lower boiling range, which overlaps that of kerosine, is charged at a controlled rate to the same tank wherein it is blended with the kerosine by a rapidly rotating agitator (not shown) which induces thorough mixing of the contents of the tank. Both of the charge stream contain substantial amounts of paraffins and aromatics, but the naphtha has a distinctly higher naphthene content than the kerosine.

The transfer -line 8 carries the mixed charge to the aromatics removal unit 10 where it is subjected to solvent extraction with sulfur dioxide. This unit also is depicted in schematic fashion since details of its operation are not germane to the present invention. The extraction solvent is introduced into the equipment in line 12, and the extract containing about 90% of the aromatic hydrocarbons in the hydrocarbon mixture is withdrawn through the bottom line 14 while the raffinate or reduced charge stock, after removal of all sulfur dioxide, leaves in pipe 16 on its way to the high pressure pump 18. This device pumps the liquid at a pressure slightly above 500 p.s.i.g. through the valved conduit 20, where the liquid meets a stream of recycle gas entering from conduit 22. This recycled material consisting predominantlyA of hydrogen along with minor amounts of methane and other gaseous hydrocarbons is charged at a rate sufficient to introduce 1000 s.c.f. of hydrogen per barrel of liquid fiowing through line 20.

The gas-liquid mixture is completely vaporized in the heater or furnace 24 in which its temperature is raised to 700 F. prior to passing through the pipe 26 en route to the catalytic pretreating unit 28 wherein its sulfur content is reduced from 750 p.p.m. to less than 5 parts per million (ppm.) in fiowing downward through a bed of cobalt molybdate-alumina catalyst of 14s-inch particle size at an average temperature of 700 F., hydrogen partial pressure of 465 p.s.i. and a liquid hourly space velocity of 3.0.

The hydrotreated material is withdrawn in conduit 30,

cooled in cooler 32 to a temperature of 125 and thenV transferred via line 33 to the stripper 34. Here the gaseous phase containing substantially all of the hydrogen sulfide and ammonia produced in hydrotreater 28 is stripped from the liquid phase by a stream of recycle gas introduced from line 31 and then withdrawn through the conduit 35 along along wtih the stripping gas. An appropriate proportion of the stripped gas mixture is vented through pipe 36 and the balance is passed through line 37 to the bottom of scrubber 38. This scrubbing unit is divided into two sections. Aqueous caustic soda is introduced into the lower section and travels downward countercurrent to the gas stream to remove hydrogen sulfide from the gas stream. In the upper section, which is provided with separate drainage for the liquid, a water wash absorbs any ammonia in the stream of gas. The purified gas phase is then carried to the main gas recycle system in valved conduit 39. The liquid phase leaves stripper 34 through the bottom line 40 where it joins another stream of the recycle gas introduced from conduit 42 at a hydrogen charging rate of 8000 s.c.f./b. of reduced, hydrotreated liquid charge stock.

In those instances where purification of the feedstock is not required, the aforesaid heating, hydrotreating, stripping and scrubbing equipment is shutdown and the reduced charge stock is bypassed from line 20 to the conduit 40 by means of the valved line 44 after the valves in lines 20, 39 and 40 are closed.

Next, the gas-liquid mixture entering pipe furnace 46- from line 40 is heated to a temperature of 870 F. and transferred through conduit 48 to the conversion reaction 50 which contains a bed of catalyst particles of 1/l-inch size. This catalyst consists of 0.35% finely dispersed platinum and 0.20% chlorine on eta alumina.

The chief conversion of the hydrocarbon charge takes place while passing downwardly through the catalyst bed in this reactor at total pressure of 450 p.s.i.g. and space Velocity of 1.5 (LHSV) with the product gases maintained at a reactor outlet temperature of 810 F. The temperature is controlled by 4suitable regulation of the firing of furnace 46. These conditions produce a dehydrogenation of about 75% of the naphthenes charged to the reactor as well as the isomerization of a considerable quantity of the parafiinic components and a minor amount of the naphthenes.

The vapor phase efiiuent from this conversion is withdrawn at the bottom of the reactor in pipe 52 and directed through. a cooler 54 and conduit S6 in being transferred tot the fractionation unit 58. The latter unit is a fractional distillation tower (schematically represented) in the instant embodiment inasmuch as it is necessary to separate three fractions here by reason of the inclusion of the low boiling blending stock naphtha through charge line 6. A normally gaseous fraction composed predominately of hydrogen plus methane and other light hydrocarbons formed by cracking side reactions in a minor degree in reactor 50 is taken overhead through conduit 60. This is the major portion of the recycle gas which is joined by a minor amount in the purified stream from pipe 39 and compressed to a pressure of 525 p.s.i.g. in recycle compressor 62 before delivery to the line 64. Most of the compressed gas is then directed into the valved line 42 as the hydrogen-rich circulating gas for the dehydrogenation reaction. A smaller amount of it is recycled via valved line 22 to supply hydrogen for the hydrogen pretreatment, and another stream of this gas is carried in conduit 31 to stripper 34 to serve as the stripping gas. Alternatively, when the hydrotreating step is omitted, an amount similar to that vented through line 36 is taken off through the valved conduit 66 for use as fuel or other purposes to avoid the build-up of inerts in the system,

A second or middle fraction of the reactor efliuent is withdrawn through line 68 for use as a motor fuel blending stock. This is a naphtha partly reformed in the relatively mild conversion and boiling below the jet fuel range, that is below about 350 F. It is desirable to separate the naphtha at this stage in order to reduce the load on -the final dearomatization unit and also to allow this normally liquid fraction to retain aromatic hydrocarbons which enhance its antiknock rating.

The third fraction which boils within the S50-550 P. range is Withdrawn through the lower line 70 'and passed through the cooler 72 and conduit 74 in transit to the aromatics removal unit 76. Aromatic hydrocarbons are removed from the kerosine conversion product here at a temperature of -10 to +20 in similar manner as in the initial dearomatization unit 10. The sulfur dioxide extractant is charged through the line 78, and the eX- tract containing about 90% of the aromatic hydrocarbons charged to the unit is removed through the bottom conduit 80; while the product, a jet fuel containing only a small percentage of aromatic hydrocarbons, is Withdrawn through the upper line 82.

A bottoms line 69 is also provided in the fractionation system for the separation of heavier hydrocarbons which boil at temperatures above the kerosine range. When a naphthene-rich gas oil (50G-600 F. boiling range) is employedr as the blending stock to be charged with a kerosine feedstock of low naphthene content,l the overhead gas and the dehydrogenated kerosine are withdrawn in pipes 60 and 70 respectively as before, and a lgas oil boiling above 550 F. is taken olf through the bottoms line 69 while the naphtha line 68 is usually shut off. Alternatively, the total charge to the plant may include as naphthene-rich blending stocks both a fraction lighter than kerosine and another fraction heavier than kerosine, plus the principal kerosine feed, in which case the conversion eiuent will be separated in the fractionator into four or more product and by-product fractions.

In another embodiment of the invention wherein the valve in line 6 is closed and no blending stock is charged to the system, a simpler separation unit 58 is employed. In this instance, the cooler 54 is operated to bring the temperature of the conversion etiiuent in conduit 52 down near ambient temperature and the unit 58 is a simple gasliquid separator. No low or high boiling liquids are taken off in lines 68 and 69; but the -gaseous phase is withdrawn to the overhead line 60, and the entire liquid phase is removed through the lower line 70 for the nal removal of aromatic compounds in the manner described hereinbefore. Using this arrangementof the apparatus, a kerosine is converted into jet engine fuel under the cornparable conditions listed hereinafter both without and with the preliminary extraction of aromatic hydrocarbons according to the present invention and Without this step.

Dehydrogenation/Isomerization:

Reaction Conditions:

Catalyst 0.35% Pt on A120;

Pressure, p.s.i.g 400 400 Space Velocity, LHSV 1. 5 1. 5 Temperature, F.3 835 810 H2 Recirculation Rate, 8, 000 8,000 H2 Partial Pressure, p.s. 340 340 Hydrogen Make s.c.f./b7 300 360 Dehydrogenated/Isomerized Product:

Boiling above 375 F.:

Yield, percent vol: y

On Process 88. 0 94. 0 On Raw Kerosine 88. 0 75.8 Component Analysis, percent vol:

En 58. 0 68. 8 1. 0 1. 0 5. 1 5. 9 Aromatics 35. 9 24. 3

Total 100.0 100. 0 Aromatics Extraction:

Process Solvent Extraction with SO2 2 Yield of J'et Fuel, vol. perent:

On Extraction 59. 7 74. 0 On Raw Kerosine. 52. 5 56.0 Properties o Product I et Fue Gravity, API 53.0 53. 0 Aniline Point, F 178.0 178. 0 Aniline-Gravity Product 9, 450 9, 450 Net Heat of Combusion, B.t.u./lb 18, 910 18, 910 Y Lumininometer Number 110 110 Freeze Point, F -40 -38 Component Analysis, percent vol Paratiins 88. 4 88. 4 O1ens 1.0 1.0 Naphthenes. 7. 6 7. 6 Aromatics 3. 0 i 3. 0

Total 100.0 100.0 Sulfur, p.p.m Nil Nil Nitrogen, p.p. Nil N11 NOTES:

1 Comparative example in first column. 2 Based on low-temperature SO2-extraction Where the extract contains vol. percent aromatics. Y

Temperature of efliuent leaving the catalyst bed.

l The tabulated data. demonstrate that the borderline production of hydrogen of 300 s.c.f./b. of charge can be increased by the prior aromatics removal -of the present process to 360 s.c.f. or more which makes the entire process, including the Vhydrotreating and stripping steps, self-sufficient in hydrogen, thereby rendering commercial operation feasiblev in a refinery having only one catalytic reformer. In addition the overall yield of jet fuel-is increased 3.5% to 56.0% (about 7% relative increase) of the volume of the raw kerosine with improvement inthe luminometer number While the severity of the conversion is diminished by reducing the reactor outlet temperature by 25 F. By reason of the reduction in the volume of material charged to the dehydrogenation reactor, the capacity of this system for producing jet fuel is increased approximately 24%. Moreover, the product analysis and properties are indicative of a jet fuel of excellent combustion characteristics and suitable for use in supersonic aircraft.

EXAMPLE2 l A l B C D Crude Source Aguaytia Wilmington Blend 100A: Blend 100A:

& Kuwait. 53B. 53B. Type o Stock Kerosine Mixed Raw Dearoma- Naphthas. tized. Boiling Range,F 30G-500 28o-400 Composition, Vol.

Percent:

Acyclics 77.7 47.5 67.3 73.8. Naphthenes 17.1 34.4 23.1 25.2. Aromatics 5.2 18.1 9.6 1.0.

The process of Example 1 is repeated under the conditions described therein, again without employing a blending step, using the Peruvian kerosine of column A of the table as the feedstock. After the initial removal of about 90% of the aromatic hydrocarbons, the charge to the dehydrogenation reaction still contains only 17.9% of naphthenes by volume and the net make of hydrogen in the dehydrogenation reaction amounts to only 200 s.c.f./b. of the reduced kerosine, which is not adequate to make up for hydrogen losses and consumption of hydrogen in the hydrogen pretreatment zone.

To overcome this hydrogen deficiency, a different feedstock is made up by blending 100 parts by volume the same kerosine with 53 parts of the naphtha mixture of column B. The original composition of the resulting blend is given in column C; and this is changed to that listed in column D by the initial aromatics extraction step. The dearomatized blend is then successively subjected to the balance of the operations described in Example 1, namely desulfurization, catalytic dehydrogenation, fractionation to remove all of the gases and liquid boiling below 350 F. from the dehydrogenation effluent and nally to further extraction of aromatic compounds from the 350-l-liquid fraction. The hydrogen make during the conversion reaction amounts to 350 s.c.f./b., and this is found to be adequate to replenish all hydrogen lost and consumed in the system. The lighter liquid fraction is utilized as a blending stock for motor gasolines, and the heavier product is an excellent jet fuel of similar composition `and properties to that set forth in Example 1.

While the invention has been described with particular reference to a few detailed examples and embodiments, it is to be understood that these have been set forth for the purpose of illustrating rather than restricting the invention. Those skilled in the art will appreciate that various modifications can be made in the specific embodiments without departing from the scope or the spirit of the invention. Accordingly, the present invention should not be construed as limited in any particular conditions or aspects except as may be set forth in the appended claims or required -by the prior art.

I claim: v

1. In producing jet engine fuels, the process which comprises removing at least a substantial proportion of the aromatic hydrocarbon content of a hydrocarbon feedstock containing parainic, naphthenic and aromatic hydrocarbons in which .the content of naphthenes is less than about by volume, hydro-treating the feedstock under desulfurization conditions in the presence of a hydrogen-rich gas and a desulfurization catalyst, thereafter` subjecting a charge containing the reduced hydrotreated feedstock with a content of more than about 25% naphthenes by volume to dehydrogenation conditions in the presence of a recycled hydrogen-rich gas and a dehydrogenation catalyst for naphthenes with the conversion severity controlled to minimize cracking and olefin formation, separating the normally gaseous and normally liquid fractions of the dehydrogenation reaction effluent, recycling a substantial portion of said gaseous fraction to the dehydrogenation and hydrotreating steps as the hydrogen-rich gas, removing at least a substantial portion of the aromatic hydrocarbons in said liquid fraction and withdrawing the remainder of said liquid fraction as a product of the process, whereby the dehydrogenation reaction produces hydrogen at a rate at least equivalent to the requirements of said process.

2. A process according to claim 1 in which said dehydrogenation catalyst is a nely divided noble metal supported on a particle form carrier.

3. A process according to claim 1 in which said hydro-` carbon feedstock is a kerosine boiling within the range of about 300 to 600 F.

4. A process according t0 claim 1 in which a kerosine tion of said charge, and said dehydrogenation reaction eiiiuent is fractionated into a normally gaseous fraction,4 a dehydrogenated kerosine and at least one liquid frac-V tion boiling outside the kerosine range prior to the re-` moval of aromatic hydrocarbons from said dehydrogenated kerosine in preparing a jet fuel.

6. A process according to claim 1 in which a kerosine of relatively low naphthene content is blended prior to said dehydrogenation with a naphtha having a substantially higher naphthene content in the preparation of said charge, and said dehydrogenation reaction eluent is fractionated into a normally gaseous fraction, a partly re.

formed hydrocarbon fraction boiling below about 350 F. and a dehydrogenated kerosene prior to the removal of aromatic hydrocarbons from said dehydrogenated kerosine in preparing a jet fuel.

7. In producing jet engine fuels, the process which comprises removing a major proportion of the aromatic hydrocarbon content of a Ikerosine containing parainic, naphthenic and aromatic hydrocarbons, hydrotreating the kerosine under desulfurization conditions in the presence of a hydrogen-rich gas and a desulfurization catalyst, therafter generating at least 340 s.c.f. of hydrogen per barrel of a charge containing the reduced hydrotreated kerosine under dehydrogenation conditions in the presence of a recycled hydrogen-rich gas and a dehydrogenation catalyst for naphthenes with the conversion severity controlled to minimize cracking and olefin formation, separating the normally gaseous and normally liquid fractions of the reaction eluent, recycling a substantial portion of said gaseous fraction to the dehydrogenation and hydrotreating steps as the hydrogen-rich gas, removing at least a substantial portion of the aromatic hydrocarbons in said liquid fraction and withdrawing the remainder of said liquid fraction as a product of the process, whereby the dehydrogenation reaction produces hydrogen at a rate at least equivalent to the requirements of said process.

8. In the production of jet engine fuels, the steps which comprise removing a major proportion of the aromatic hydrocarbon content of a kerosine boiling within the range of about 300 to 600 F. and containing parafns, aromatic hydrocarbons and less than about 25% by volurne of naphthenes, and hydrotreating the kerosine under desulfurization conditions in the presence of a recycled hydrogen-rich gas and a desulfurization catalyst to produce a reduced hydrotreated kerosine containing more than about 25% of naphthenes, thereafter subjecting said reduced hydrotreated kerosine to dehydrogenation-isomerization conditions in the presence of a recycled hydrogen-rich gas and a noble metal catalyst for the dehydro genation of naphthenes and isomerization of parafl'ns with the conversion severity controlled to minimize crack ing and olen formation, separating a normally liquid fraction of the reaction effluent from a hydrogen-rich normally gaseous fraction, recycling a minor portion of said gaseous fraction to said hydrotreating step, recycling a major portion of said gaseous fraction to said dehydrogenation-isomerization step, removing a major proportion of the aromatic hydrocarbons in said liquid fraction and withdrawing the resulting reduced liquid fraction as a product of the process, whereby the dehydrogenation reaction supplies all of the hydrogen required in said process.

9. A process according to claim 8 in which said dehydrogenation-isomerization catalyst contains between about 0.1 and 1.0% platinum supported on particle form alumina.

10. A process according to claim 8 in which said dehydrogenation-isomerization conditions include a hydrogen partial pressure between about 20 and 1400 p.s.i., a reaction eluent temperature between about 640 and 980 F., a hydrogen charging rate between about 1000 and 15,000 s.c.f./b. of reduced hydrotreated kerosine and a liquid hourly volumetric space velocity between about 0.5 and 60.

11. A process according to claim 8 in which said dehydrogenation-isomerization conditions include a hydrogen partial pressure between about 100 and 500 p.s.i., a reaction eluent temperature between about 780 and 850 F., a hydrogen charging rate betwen about 4000 and 10,000 s.c.f./ b. of reduced hydrotreated kerosine, a liquid hourly volumetric space velocity between about 1 and 20 and between about 0.1 and 1.0% by weight of a noble metal catalyst supported on an inert particle form carrier.

12. A process for the production of jet fuels which comprises blending a kerosine boiling within the range of about 350 to 550 F. and containing parains, aromatic hydrocarbons, and a relatively loW content of naphthenes with a hydrocarbon stock having a substantially higher naphthene content, removing a major proportion of the aromatic hydrocarbon content of the resulting blend, hydrotreating the reduced blend underdesulfurization conditions in the presence of a recycled hydrogen-rich gas and a desulfurization catalyst, thereafter subjecting the reduced hydrotreated blend to dehydrogenation conditions in the presence 0f a recycled hydrogen-rich gas and a catalyst for the dehydrogenation of naphthenes with the conversion severity controlled to minimize cracking and olen formation, fractionating the reaction effluent into a hydrogen-rich normally gaseous fraction and a liquid fraction boiling within the range of about 350 to 550 F., recycling a minor portion of said gaseous fraction to said hydrotreating step, recycling a major portion of said gaseous fraction to said dehydrogenation reaction, removing a major proportion of the aromatic hydrocarbons in said liquid fraction and withdrawing the resulting reduced liquid fraction as a jet fuel, whereby the dehydrogenation reaction supplies all of the hydrogen required in said process.

References Cited UNITED STATES PATENTS 3,030,299 4/ 1962 Plummer 208-96v 3,110,661 11/1963 Franz 208-96 3,201,342 8/1965 Bachman et al. 208-89 3,230,165 1/ 1966 Cunningham 208-89 DELBERT E. GANTZ, Primary Examiner.

S. P. JONES, Assistant Examiner.

Claims (1)

1. IN PRODUCING JET ENGINE FUELS, THE PROCESS WHICH COMPRISES REMOVING AT LEAST A SUBSTANTIAL PROPORTION OF THE AROMATIC HYDROCARBON CONTENT OF A HYDROCARBON FEEDSTOCK CONTAINING PARAFFINIC, NAPHTHENIC AND AROMATIC HYDROCARBONS IN WHICH THE CONTENT OF NAPHTHENES IS LESS THAN ABOUT 25% BY VOLUME, HYDRO-TREATING THE FEEDSTOCK UNDER DESULFURIZATION CONDITIONS IN THE PRESENCE OF A HYDROGEN-RICH GAS AND A DESULFURIZATION CATALYST, THEREAFTER SUBJECTING A CHARGE CONTAINING THE REDUCED HYDROTREATED FEEDSTOCK WITH A CONTENT OF MORE THAN ABOUT 25% NAPHTHENES BY VOLUME TO DEHYDRATION CONDITIONS IN THE PRESENCE OF A RECYCLED HYDROGE-RICH GAS AND A DEHYDROGENATION CATALYST FOR NAPHTHENES WITH THE CONVERSION SEVERITY CONTROLLED TO MINIMIZE CRACKING AND OLEFIN FORMATION, SEPARATING THE NORMALLY GASEOUS AND

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