US2773008A - Hydrofining-hydroforming system - Google Patents

Description

R. J. HENGSTEBECK 2,773,008

HYDROFINING HYDROFORMING .SYSTEM Y Dec. 4, 1956 Filed Apr .1 26; 1954 DIL 1mm mY IN V EN TOR Robert J. Hengslebeck @7@ ATTORNEY SS mw\ 1 m www R\E United States Patent' HYDROFINING-HYDROFORMING SYSTEM Robert J. Hengstebeck, Valparaiso, Ind., assignor to ntandard Oil Company, Chicago, Ill., `a corporation of diana Application April 26, 1954, Serial No. 425,461

12 Claims. (Cl. 196-28) This invention relates to an improved hydrofininghydroforming system wherein hydrogen produced in a hydroforming operation is employed for effecting hydrofining of a plurality of stocks such as naphtha and/ or one or more distillate fuel oils.

Various systems have been proposed for combining hydrofining and hydroforming operations whereby the hydrogen produced by hydroforming is utilized in hydrofining but all of such systems have required large capital investment and operating costs. An object of this invention is to provide a new and improved unitary hydrofining hydroforming combination of remarkably low investment and operating expense. A further object is to decrease compression costs and to decrease the number of required compressors. An important object is to eliminate substantially recycle of low boiling hydrocarbons in the hydroforming and hydrofining reactors without requiring the necessity of expensive fractionation systems. Other objects will be apparent as the detailed description of the invention proceeds. l

In the preferred application of the invention the hydro forming is effected with a platinum-on-alumina catalyst by the so-called Ultraforming process as'descbed by W. I. Birmingham in The Petroleum Engineer, volume XXVI (April 1954), page C-35, and in co-pending application Serial No. 347,635, filed April 9, 1953. The cata- It' is desirable that virgin distillate fuelv oils be kept separate from cracked distillate fuel oils and, n the preferred embodiment of the invention, such streams are separately hydrofined. The hydrogen stream from the first hydrotining effluent separator is heated with a virgin distillate fuel and recycled hydrogen and hydroflning is effected in a second hydrofining zone at a slightly lower pressure and at a temperature of about 650 to 800, e. g. 775 F. The separator following the second hydrofining step is likewise operated at substantially hydrofining pressure and at a temperature in the range of about 200 to 500 F. so that, here again, low boiling hydrocarbons such as butanesand pentanes remain uncondensed and leave the separator with the hydrogen stream. 'I'he low boiling constituents of the condensed liquid are removed by fractionation to obtain a distillate fuel of desired boiling range and liash point and the low boiling components may be combined with desulfurized naphtha to form a part of p the hydroformer charge.

The hydrogen stream, still containing HzS and light condensable hydrocarbons from the first and second hydrofiningefiluent separators, is vthen combined with a cracked lyst employed for hydroning (hydrodesulfurizaton) is y l preferably of the type known as sulfur immune hydrogenation catalysts, i. e., a combination of group VI metal oxide or sulfide with a group VIII metal oxide or sulfide supported on alumina gel. The net hydrogen produced in the hydroforming step is employed for hydrofining the naphtha which is to be charged to the hydroformer and the I-IzS-containing hydrogen stream separated from the first hydrofining product is subsequently utilized without further compression for hydrofining a virgin distillate fuel oil or a cracked distillate fuel oil or each of said distillate fuel oils in sequence without further' recompression of hydrogen.

The naphtha charge is hydrofined at 200-1500, preferably about 750 p. s. i., at a temperature of about 600800 F. with a hydrogen stream which contains some condensable hydrocarbons such as butanes in a first hydr'oiining or hydrodesulfurization step. A hydrogen separator following the first hydrofining step is operated at substantially hydrofining pressure, i. e. 200 to 1500, e.,g. about 750 p. s. i. g., and at a temperature in the range of about 200 to 500 F. so that substantiallyV all of the butanes' and pentanes leave the separator with the hydrogen stream and the condensate from the separator maybe directly charged to the hydroformer without fractionation. Operation of this separator at the defined high temperature and pressure, preferably with hydrogen stripping, effects substantial savings in operating costs since less reheating of the separated streams Ais required for their subsequent use and since HzS and light' hydrocarbons are removed more completely from the condensedhydrocarbonsl at this relatively high temperature level.

distillate fuel such as light catalytic cycle oil which is hydrofined in a third hydrofining reactor at a pressure somewhat below that of thesecond hydrofining reactor and at a temperature in the range of about 650 to 800 F. The efiiuent leaving the third hydrofining reactor is cooled to as low a temperature as can be readily obtained with available cooling water, e. g. 100 F., or lower so that in this case the pentanes and butanes are largely condensed and separated by fractionation as a gasoline or gasoline blending stock from a distillate fuel oil of desired boiling range and flash point. A part of the hydrogen from the final separator is vented from the system, preferably through an adsorber to pick up any condensable hydrocarbons contained therein. Most of the hydrogen stream, however, is passed through a scrubber for removing hydrogen sulfide and the purified hydrogen is recycled to one or more of the hydrofining steps, a part of the purified hydrogen being employed for stripping HzS and'condensable hydrocarbons out of condensate which is separated at ZOO-500 F.

The invention will be more clearly understood from the following description of a commercial unit for hydrofining a hydroformer charge consisting of 3,300 barrels per day of virgin naphtha and 2,030 barrels per day of coke still naphtha, separately hydrofining 2,500 barrels per day of a virgin fuel oil and finally hydrotining 3,860 barrels per day of light catalytic cycle oil. This example is illustrated in the accompanying drawing which forms a part of this specification and which is a schematic flow diagram of the hydrofning-hydroforming system.

In this example the invention is applied to a refinery in which a crude is fractionated into a virgin naphtha of 56 A. P. I. gravity and boiling in the range of about to 360 F., a virgin distillate fuel fraction of about 40 A. P. I. gravity boiling from about 360 to 500 F. (No. l fuel oil), a gas oil which is subjected to catalytic cracking and a residue which is subjected to coking. A light cycle gas oil from the catalytic cracking operation having an A. P. I. gravity of about 24, boiling range of about 430 to 600 F. and a sulfur content of about 2.4 weight percent is withdrawn from the catalytic cracking operation as a No. 4 fuel oil stock, the heavy catalytic cycle oil being charged to the coker. The coke still naphtha having an A. P. I. gravity of about 55, a boiling range of about 175 to 400 F. and a sulfur content of about 1.1 percent is combined with the virgin naphtha, and the coke still gas oil is charged to the catalytic cracking operation. The mixture of virgin ,and coke still naphtha "is hydrofined (hydrodesulfurized) and then hydroforrned, preferably by the Ultraforming process,

to give high octane number gasoline or blending stock. The virgin distillate fuel is hydrofined to give a high quality No. l fuel oil having a tag closed cup flash point of at least 130F., a sulfur content of not more than .1 percent, a color of +14 Saybolt and superior burning quality. The light catalytic cycle oil is hydroned to give a No. 4 fuel oil having a Pensky-Martens flash point of at least 140 F., a viscosity of about 33-40 Saybolt Universal seconds at 100 F., a maximum A.V S. T. M. 10 percent point of 440 F., a sulfur content of not more than .7 percent and a pour point of 5 F. While coke still naphtha is employed as a part of the first hydrofning charge in this example, it should be understood that the invention is applicable to theV use of any high sulfur cracked naphtha.

The virgin naphtha from line and coke still naphth from line 11 are introduced by pump 12 at about 780 p. s. i. g. together with hydrogen from line 13 through preheater 14 to rst hydroning reactor 15. The hydrogen from line 13 is a net hydrogen stream from the hydro-l forming step, it may contain about 80 mol. percent hydrogen and the balance C1-C5 light hydrocarbons and it may be added tothe mixed naphtha charge at the rate of about 800 to- 1000 cubic feet (standard conditions) per barrel. The charge and hydrogen are preheated to a temperature of about 625 to 700 F. and further heated by exothermic heat of conversion to a temperature in the range of about 700 to 800 F. Reactor 15 contains about 625 cubic feet of hydroiining catalyst having a pill size of about 3/16 inch and a bulk density of about 60 pounds per cubic foot. The catalyst is preferably an aluminav support containing about 3 percent cobalt 'as cobalt oxide and about 9 percent molybdenum as molybdenum oxide and it may be prepared by impregnating an alumina gel catalyst support with ammonium molybdate and then, either before or after drying and calcining, with cobalt nitrate followed by conventional drying and ealcining.V Any conventional hydroning catalyst may be employed but this catalyst is preferably of the so-called sulfur immune type, i. e. a mixture of the oxides or sultides of a group VI metal such as' molybdenum, chromium or tungsten and a group VIII metal such as cobalt or nickel on an alumina support, the so-called cobalt molybdate or cobalt-molybdena-on-alumina catalysts being preferred. In the first hydroning reactor there is hydrogen consumption both by the virgin naphtha and the coke still naphtha components of the feed although the latter accounts for most of the hydrogen consumption.

The effluent stream from reactor 15 passes through line 16 and heat exchanger 17 to high pressure separator 18 which is operated at substantially reaction pressure but at a temperature in the range of about 200 to 500 F., about 400 F. in this example, so that most of the butanes and pentanes and some of the hexanes remain uncondensed and are withdrawn with the HzS-containing hydrogen stream through line 19, the condensed naphtha boiling in the approximate range of about 170 to 380 F. and containing not more than about .006 weight percent sulfur being withdrawn by line 20 for charging to the hydroformer. The temperature at which the separator is operated will be dependent upon the pressure, lower temperatures being employed for lower pressures. Only a rough'separation is effected between light ends (C5 and lighter hydrocarbons) and heavy ends (C7 and heavier hydrocarbons) some of the Ce hydrocarbons being condensed and some remaining uncondensed. In order to insure removal of substantially all HgS and light ends from the condensate, separator 18 is preferably a vertical stripper vessel. Substantially HzS-free hydrogen is introduced at the base of the stripper through line 23a for stripping H28 and light ends out of the condensate.

The virgin distillate fuel is introduced by pump 21 and line 22 at a pressure of about 760 `p. s. i., ,admixed with the hydrogen stream from line 19 and that portion v(it avr-apos ,r r

any) of the recycled hydrogen stream from line 23 which was not used as stripping gas in separator or stripping tower 1S, and the mixed stream is passed through heater 24 to hydrotining reactor 25 with an inlet temperature of about 700 to 750 F. so that the maximum temperature in reactor 25 will be about 700 to 800 or about 750 F. in this reactor the catalyst is the same as in reactor 15 but the amount of catalyst is only about 120 cubic feet. Here again hydrogen is consumed and there is additional production of some normally gaseous hydrocarbons.

The effluent stream from reactor 25 passes by line 26 through heat exchanger 27 to high pressure separator 28 which, like separator 18, is operated at substantially reaction pressure and at a temperature in the range of about 200 to 500 F., e. g. about/100 F., so that the Ci--Cs light hydrocarbons remain uncondensed and leave the separator with the hydrogen stream through line 29. Here again a stripping tower may be used for effecting more complete separation, the stripping gas being hydrogen introduced through a line (not shown) from line 23 to the base of the stripping tower. The condensate from separator 28 is introduced by line 30 to a fractionator diagrammatically illustrated by tower 31 provided with heater 32. Any remaining light hydrocarbons and gases may be withdrawn through line 33 and naphtha boilingrange hydrocarbons are introduced by line 34 to line 20, the fractionation system being operated to produce a high quality distillate fuel oil of the desired boiling range and flash point, which fuel oil is withdrawn through line 35.

The naphtha from line 20, and preferably also from line 34,.is combined with about 4,000 to 6,000 or more cubic feet perrbarrel of hydrogen from line 36, passed through preheater 37 and introduced into the rst reactor 38 of the hydroforming system which, in this example, is an Ultraforrner. A pressure reduction valve 20a may be employed to reduce the pressure of the mixed naphtha charge to about 340 p. s. i.; if desired, a separatoror stripper (not shown) may be employed in the line immediately following the pressure reduction in order to remove any released HzS but the use of such separator or fractionating means is usually not necessary because of the high temperature at which separators 18 and 28 are operated and the stripping effected therein. The catalyst-in reactor 38 yis of the platinum-on-alumina type preferably containingabout .3 to .6 weight percent platinum and prepared, for example, as described in U. S. Any known type of platinum-on-alumina hydroforming catalyst may be employed and the catalyst may contain about .l to 1 percent of a halogen such as chlorine or iluorine. Such catalysts are well known to those skilled in the art and require no further description. The catalyst inv this example has a pill size of about 1/s inch, a bulk density of about 50 pounds per cubic foot and each of the hydroformer reactorsV contains about cubic feet of such catalyst.

The mixed naphtha and hydrogen stream enters reactor 38 at a temperature of about 950 F. but is cooled by the exothermic hydroforming reactions. The eiuent from the rst reactor passes by line 39 to reheater 40 and thence to second reactor 41 again with an inlet temperature of about 950 F. and the efuent from the second reactor passes by line 42 through reheater 43 and thence to third reactor 44 at a temperature of about 950 F. Each of the hydroforming reactors contains the same amount of platinum-on-alumina catalyst and, although not shown in the drawing, a fourth reactor is usually employed as an alternate for the on-stream reactors so that the catalyst in any one of the reactors may be replaced by the alternate reactor and subjected to regeneration and rejuvenation. As far as this invention is concerned, a non-regenerative hydroforming system such as convenv tional Platforming may be employed.

The e'u'ent from the nal hydroforming reactor passes by line 45 through heat exchanger 46 and cooler 47 to separator 48 which is operated at substantially hydroforming pressure and at a temperature obtainable by available cooling water, usually not higher than about 100 F. Most of this hydrogen stream is recycled by compressor 50 through line 36 and a heat exchanger (e. g. 46) for use in the hydroforming step but the net hydrogen stream is compressed by compressor 51 for introduction through line 13 to the entering mixed naphtha charge as hereinabove described. The hydroformed naphtha is Withdrawn from the separator through line 52 and depropanized, etc., in accordance with conventional practice.`

The light catalytic cycle oil is introduced by pump 53 through line 54 at approximately 750 p. s. i. and admixed with the hydrogen stream from line 29 which, as above described, contains not only HzS but C1--C5 light hydrocarbons as well, Ithe mixture passing through heater 55 to hydrofining reactor 56 at an inlet temperature of about 600 to 700 F. so that the maximum temperature produced in reactor 56 is in the range of about 700-800" F. Reactor 56 contains about 180 cubic feet of the same type of catalyst employed in reactors 15 and 25, respectively. The effluent stream from reactor 56` passes through line 57 and heat exchanger 58 to high pressure temperature as can be obtained with available cooling Water, i. e. about 100 F., or lower. The uncondensed hydrogen and HzS leave separator 59 through line 60 and a net tail gas is vented through line 61, preferably through an absorber for recovering condensable hydrocarbons. A portion of the hydrogen, however, is compressed by compressor 62 and passed through an amine type scrubber system 63 for eliminating HzS through line 64, the purified hydrogen being recycled by line 23 for stripping condensate, admixture with the virgin fuel oil charging stock introduced by line 22 as hereinabove described, or introduction to other of the hydrofining reactors.

The condensed hydrocarbons from separator 59 are passed by line 65 to a fractionation system diagrammatically represented by tower 66 containing heater 67, the C3 and lighter gases being discharged by line 68, a gasoline fraction by line 69 and the finished No. 4 fuel oil through line 70.

From the foregoing description it will be observed that the objects of the invention have been attained. The hydrogen streams from separator 18 and from separator 28, respectively, require no compression since reactor 25 is operated at sufficiently lower pressure than reactor 15 and Vreactor 56 is operated atsufficiently lower pressure than reactor 25 so as to avoid the necessity of recompressing the hydrogen which passes downstream in this circuit. The operation of separators 1S and 28 at relatively high temperatures in the range of 200 to 500 F. minimizes the cooling thatV must be effected by exchangers 17 and 27,` minimizes the amount of heat that must be supplied by heaters 24, 32, 37, and 55 and, coupled with the hydrogen stripping, avoids the build-up of light ends in the hydroformer without requiring expensive fractionation equipment. The hydrogen employed in reactor 15 is preferably on a once-through basis so that no H2S separation step is required at this point. A single HzS scrubber is employed for removing HzS from hydrogen recycled in the distillate fuel hydrofining portion of the system but this scrubber can be relatively small. The light ends, i. e. C4, C5 and some Ce hydrocarbons, which are present in the reactor efliuent from reactors 15, 2S and 56 are substantially all recovered as a part of the gasoline from fractionator 66 (although a minor amount may be recovered from line 33 and by scrubbing tail gas vented through line 61) The unitary system hereinabove described thus effects enormous savings in equipment and operating costs, obtains effective utilization of hydrogen,

, separator 59 which, 4in this case, is operated at as low a avoids inerts build-up in the hydroformer and hydrofining steps and produces remarkably high quality hydroformed naphtha as well as distillate fuel oils.

While a preferred `example of the invention has been described in great detail, it should be understood that the invention is not limited thereto since alternative arrangements and operating conditions willbe apparent from the above description to those skilled in the art. For example, at least a part ofthe hydrogen from scrubber 63 may be introduced along with hydrogen from line 13 to the naphtha charge entering preheater 14 and/or along with hydrogen from line 29 entering preheater 55. The naphtha may be entirely a virgin naphtha or entirely a cracked naphtha although the invention is particularly applicable to a high sulfur thermally cracked naphtha such as coke still naphtha. The hydroforming may be effected by any known catalytic naphtha reformingprocess which produces by-product hydrogen. Other types of hydrofining catalysts such as platinum-,on-aluminamay be employed. At least one of the hydrofining steps may employ residual stocks instead of distillate fuel oil stocks with the hydrofining temperature being adjusted accordingly.

I claim:

. l. A hydrocarbon conversion process which comprises hydrofining a sulfur-containing naphtha at a pressure in the range of about 200 to 1500 p. s. i. g. and a temperature in the range of 600-800 F. in a first hydrofining zone, cooling the eiiiuent from the first hydrofining zone at substantially hydrofining pressure to condense hydrocarbons higher boiling than hexane withoutcondensing appreciable amounts of pentanes and lower boiling hydrocarbons, separating in a first separation zone at a temperature above 200 F. condensed hydrocarbons from a hydrogen stream containing H28 and uncondensed hydrocarbons, contacting said hydrogen stream and a distillate hydrocarbon fuel which vis higher boiling than gasoline with a hydrofining catalyst in a second hydrofining zone at a temperature in the range of about 600 to 800 F. and at a slightly lower pressure than that maintained in the first named hydrofining step so that the hydrogen stream requires no further compression between said first and second hydrofining zones, cooling the second hydrofining reaction zone eiuent, separating condensed hydrocarbons therefrom atsubstantially hydrofining pressure in a second separation zone and distilling said separated hydrocarbons to obtain 'a distillate fuel oil and a naphtha fraction, contacting said first named condensed hydrocarbons with a platinum-onalumina catalyst at a temperature in the range of about 850 to 950 F. under a pressure in the range of about 200 to 750 p. s. i. to effect hydroforming, separating a hydrogen stream containing low boiling hydrocarbons from hydrocarbons produced in the hydroforming step, recycling most of the last named separated hydrogen stream to the hydroforming step and introducing the remainder to the first named hydrofining step.

2. The process of claim 1 wherein the first separation zone is operated at a temperature in the range of about 200 to 500 F. andwherein the condensed hydrocarbons are stripped with hydrogen before being charged to the hydroforming step.

3. The process of claiml wherein the naphthafraction separated from the distillate fuel oil is also charged to the hydroforming step.

4. The process of claim 1 which includes the further steps `of removing a hydrogen stream containing HzS and low boiling hydrocarbons from the second separation zone at a temperature above 200 F. and contacting it together with recycled hydrogen and arcracked fuel oil distillate hydrocarbon with a hydrofining catalyst in a last hydrofining zone at a temperature of about 600 to 800 F. at a slightly lower pressure than that maintained in the second named hydrofining zone so that the hydrogen stream from the separation step requires no further compression, cooling the eiuent fronrthe lastv 5. The method of separately hydrogenating a naphtha,

a virgin distillate fuel and a cracked distillate fuel with by-product hydrogen produced by the hydroforming of naphtha, which method comprises compressing said byproduct hydrogen to a pressure of at least about 750 p. s. i. g., heating the compressed hydrogen and naphtha at said pressure to a temperature in the range of about 600 to 750 F., contacting the heated mixture with a Vhydroining catalyst in a first hydroning zone to etect hydrodesulfurizat-ion of the naphtha, cooling the effluent product stream to a temperature in the range of about 200 to 500 F. withoutA substantial reduction in pressure to effect condensation of C7 and heavier hydrocarbons without condensing appreciable amounts of C and lighter hydrocarbons, separating a hydrogen stream containing uncondensed hydrocarbons from condensate, heating said hydrogen stream and said virgin distillate fuel to a temperature in the range of about 700 to 750 F. at a pressure slightly lower than that maintained in the separation zone so that the hydrogen stream requires no further compression, contacting the heated stream with a hydroning catalyst in a second hydroning zone at a temperature in the range of 700 to 800 F., cooling the ei'liuent from the second hydioiining zone to a tem-V perature in the range of 200 to 500 F. without substantial reduction in pressure to condense Cv and heavier hydrocarbons without condensing substantial amounts of C5 and lighter hydrocarbons, separating a second hydroture not substantially higher than 100 F. for condensl' ing most of Vthe condensable hydrocarbons, separating a hydrogen stream from the third separation zone, vent-V ing a part of said last named hydrogen stream, scrubbing H28 from another part of said last named hydrogen stream and recycling the scrubbed hydrogen for admixture with virgin distillate fuel entering the second hydroning zone.

.6. The method of claim 5 wherein at least a part of the scrubbed hydrogen isemployed for stripping HzS and light hydrocarbons from the first named condensate before it is admixed with virgin distillate fuel.

7. The method of claim 5 which includes the steps of fractionating condensate from the second separating zone to obtain a distillate fuel and a naphtha fraction, combining said naphtha fraction with condensate removed from the first separation zone, hydroforming the mixed naphtha in the presence of hydrogen with a platinum-on-alumina catalyst, separating a hydrogen stream from the hydroforming eiiiuent, recycling the major part of the last named hydrogen stream and employing a minor part of said last named hydrogen stream for introduction with naphtha into the rst hydrotining zone.

8. The method of claim 5 which includes the step of fractionating the condensate from the third separating zone into a gasoline fraction and a distillate fuel fraction.

9. The process of claim 1 wherein the sulfur-containing naphtha comprises cracked naphtha.

10. A hydrocarbon conversion process which comprises hydroiining a sulfur-containing naphtha at a pressure in the range of about k200 to 1500 p. s. i. g. and a temperature in the range of about 600 lto 800 F. in a rst hydroning Zone, cooling the eluent from the rst hydroning Zone at substantially hydroning pressure to condense hydrocarbons 'higher boiling than hexane without condensing appreciable amounts of pentanes and lower boiling hydrocarbons, separating in a first separation zone at a temperature above 200- F. condensed hydrocarbons from a hydrogen stream containing H28 and uncondensed hydrocarbons, contacting said hydrogen stream and a hydrocarbon which is higher boiling than gasoline with a hydroning catalyst in a second hydroning zone at a temperature in the range of about 600 to 800 F. and at a hydrofining pressure slightly lower than that maintainedV in the first ynamed hydroning step so that the hydrogen stream requires no further compression between saidV hydrofining zones, cooling the second hydrotining reaction zone eluent, separating condensed hydrocarbons therefrom at substantially hydroiining pressure in Va second separation zone, distilling hydrocarbons separated from the second separation zone to remove light hydrocarbons from a heavy hydrocarbon fraction, stripping the tirst named condensed hydrocarbons to effect increased removal of HzS and light hydrocarbons therefrom, contacting said stripped lhydrocarbons with a platinum-on-alumina catalyst at a temperature in the range of about 850 to 950 F. under a pressure in the range of about 200 to 750 p. s. i. to eiect hydroforming, separating a hydrogen stream containing low boiling hydrocarbons from hydrocarbons produced in the hydroforming step, recycling most of the last named separated hydrogen stream to the hydroforming step and introducing net hydrogen produced in the hydroforming step to the first named hydroining step.

1l. The method of claim 10 wherein said stripping is effected with hydrogen. v

l2. The method of claim l0 which includes the step of introducing recycled hydrogen to said second hydrotining reaction zone.

References Cited in the tile of this patent UNITED-STATES PATENTS 2,417,308 Lee Mar. 1l, 1947 2,516,877' Horne et al. Aug. l, 1950 2,587,987 Franklin Mar. 4, 1952 2,67l,754 De Rosset et al. Mar. 9, 1954 2,691,623 Hartley Oct. l2, 1954 t l l

Claims (1)

1. A HYDROCARBON CONVERSION PROCESS WHICH COMPRISES HYDROFINING A SULFUR-CONTAINING NAPHTHA AT A PRESSURE IN THE RANGE OF ABOUT 200 TO 1500 P.S.I.G. AND A TEMPERATURE IN THE RANGE OF 600-800* F. IN A FIRST HYDROFINING ZONE, COOLING THE EFFLUENT FROM THE FIRST HYDROFINING ZONE AT SUBSTANTIALLY HYDROFINING PRESSURE TO CONDENSE HYDROCARBONS HIGHER BOILING THAN HEXANE WITHOUT CONDENSING APPRECIABLE AMOUNTS OF PENTANES AND LOWER BOILING HYDROCARBONS, SEPARATING IN A FIRST SEPARATION ZONE AT A TEMPERATURE ABOVE 200* F. CONDENSED HYDROCARBONS FROM A HYDROGEN STREAM CONTAINING H2S AND UNCONDENSED HYDROCARBONS, CONTACTING SAID HYDROGEN STREAM AND A DISTILLATE HYDROCARBON FUEL WHICH IS HIGHER BOILING THAN GASOLINE WITH A HYDROFINING CATALYST IN A SECOND HYDROFINING ZONE AT A TEMPERATURE IN THE RANGE OF ABOUT 600 TO 800* F. AND AT A SLIGHTLY LOWER PRESSURE THAN THAT MAINTAINED IN THE FIRST NAMED HYDROFINING STEP SO THAT THE HYDROGEN STREAM REQUIRES NO FURTHER COMPRESSION BETWEEN SAID FIRST AND SECOND HYDROFINING ZONES, COOLING THE SECOND HYDROFINING REACTION ZONE EFFLUENT, SEPARATING CONDENSED HYDROCARBONS THEREFROM AT SUBSTANTIALLY HYROFINING PRESSURE IN A SECOND SEPARATION ZONE AND DISTILLING SAID SEPARATED HYDROCARBONS TO OBTAIN A DISTILLATE FUEL OIL AND A NAPHTHA FRACTION, CONTACTING SAID FIRST NAMED CONDENSED HDYROCARBONS WITH A PLATINUM-ONALUMINA CATALYST AT A TEMPERATURE IN THE RANGE OF ABOUT 850 TO 950* F. UNDER A PRESSURE IN THE RANGE OF ABOUT 200 TO 750 P.S.I. TO EFFECT HYDROFORMING, SEPARATING A HYDROGEN STREAM CONTAINING LOW BOILING HYDROCARBONS FROM HYDROCARBONS PRODUCED IN THE HYDROFORMING STEP, RECYCLING MOST OF THE LAST NAMED SEPARATED HYDROGEN STREAM TO THE HYDROFORMING STEP AND INTRODUCING THE REMAINDER TO THE FIRST NAMED HYDROFINING STEP.

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