US2691623A

US2691623A - Hydrocarbon conversion process

Info

Publication number: US2691623A
Application number: US190490A
Authority: US
Inventors: Fred L Hartley
Original assignee: Union Oil Company of California
Current assignee: Union Oil Company of California
Priority date: 1950-10-17
Filing date: 1950-10-17
Publication date: 1954-10-12
Anticipated expiration: 1971-10-12

Description

Oct. 12, 1954 F, HARTLEY l HYDROCARBON CONVERSION PROCESS Filed Oct. 17, 1950 a. l wn Rmx@ AWM/me. .yf/250 Z- Awww;

Patented Oct. 12, 1954 2,691,623 HYDROCARBON CONVERSION PROCESS Fred L. Hartley, Palos Verdes Estates, Calif., as-

signor to Union Oil Company of California, Los

Angeles, Calif., a corporation of California Application October ll, 1950, Serial No. 190,490

(Cl. IBS-24) 17 Claims. l

This invention relates generally to the rening of crude petroleum and residua and more particularly relates to a new and improved process for rening crude petroleum wherein an external source of hydrogen has been eliminated.

The upgrading of petroleum stocks by catalytic processing in the presence of hydrogen has been known for many years but its effectiveness has been somewhat limited by the relatively high cost of hydrogen production. While new methods for producing hydrogen have greatly decreased the cost in recent years, this problem is nonetheless still present.

The process of this invention makes use of twof important principles in order to eliminate the use of externally supplied hydrogen. In the iirst instance crude petroleum normally having a high carbon-hydrogen ratio is rst subjected to high temperature coking in order to produce a cracked distillate having a lower carbonhydrogen ratio and which is at the same time more amenable to upgrading by hydrogenation. In the second instance the process employs straight run distillates of the gasoline boiling range, such as are normally available in fair quantity in a modern refinery, in order to adjust the hydrogen requirements. This is accomplished in a reforming step wherein the coker distillate, following a desulfurization step, is reformed simultaneously with the straight run fraction. The flow of straight run to the reforming step is regulated and controlled by the hydrogen requirements of the preceding desulfurization stage.

It is an object of this invention to refine crude petroleum and to obtain a low boiling petroleum fractions and with a minimum yield of less desirable products.

It is another object of this invention to renne crude petroleum by means of catalytic desulfurization and hydrogenation wherein no external source of hydrogen is required.

It is another object of this invention to simultaneously reform straight run and desulfurized coker distillate fractions whereby a hydrogenrich stream is obtained which supplies the hydrogen requirements of the preceding desulfurization stage.

It is another object of this invention to simultaneously renne crude oil and upgrade straight run gasoline.

It is another object of this invention to dilute desulfurized coker distillate fractions with a straight run stock in order to dilute the nitrogen content of the desulfurized fraction and thereby minimize nitrogen poisoning during a subsequent reforming step.

Briefly, this invention relates to a three-step process in the refining of crude oil, namely the sequence of coking, desulfurization and reformmaximum yield of` ing. IThe coking step is any conventional process such as delayed coking for the conversion of the oil into gases, vapors boiling below a given end point, and solid coke. In an optional step of the process the coker distillate may be subjected to additional cracking such as in a thermal or catalytic cracker in order to reduce the boiling range of the stock. The preceding coking step serves primarily 'to reduce the carbon-hydrogen ration of the fractions to be subjected to subsequent processing. The coker distillate, together with any cracked derivatives, is subjected to desuliurization step in the presence of a cobalt molybdate catalyst whereby hydrogenation of olefins, sulfur removal and'nitrogen removal are simultaneously effected. vIn the preferred modification the desulfurization step is effected at such temperatures as to minimize the synthesis of aromatics, which synthesis is accompanied by increased carbon deposition. The desulfurized product is iirst mixed with a straight run gasoline fraction and is then subjected to reforming in the presence of a platinum-alumina catalyst. The reforming step is of such character that extensive dehydrogenation and aromatization of the charge stock occurs with consequent hydro-- gen production. The hydrogen produced in the reforming step is mixed with the unreacted hydrogen of the desulfurization step and is subjected to separation by hypersorption, for example, in order to produce a hydrogen-rich recycle stream which in part flows as recycle hydrogen to the reforming step, and which in part iiows as recycle hydrogen to the desulfurization step. Such reforming in the presence of hydrogen is also referred to as hydroforming, and such desulfurization in the presence of hydrogen is also referred to as "hydro-desulfurization.

The cobalt molybdate catalyst which has been found to be effective in the process of this invention comprises a mixture of cobalt and molybdenum oxides wherein the molecular ratio of CoO to M003 is between about 0.4 and 5.0. This catalyst may be employed in unsupported form or, alternatively, it may bedistend'ed on a suitable carrier such as alumina, silica, zirconia, thoria, magnesia, magnesium hydroxide, titania or any combination thereof. Of the foregoing carriers it has been found that the preferred carrier material Ais alumina and especially alumina containing about 3-8% by weight of silica.

In the preparation of the unsupported cobalt molybdate catalyst the catalyst can be coprecipitated by mixing aqueous solutions of, for example, cobalt nitrate and ammonium molybdate, whereby a precipitate is formed. The precipitate is filtered, washed, dried and finally activated by heating to about 500 C. Alternatively, the cobalt molybdate may be supported on alumina by coprecipitating a mixture of cobalt, aluminum and molybdenum oxides. A suitable hydrogel of the three components can be prepared by adding an ammoniacal ammonium molybdate solution to an aqueoussolution of cobalt and aluminum nitrates. The precipitate which results is washed, dried and activated. In still another method a washed aluminum hydrogel is suspended in an aqueous solution of cobalt nitrate and an ammoniacal solution of ammonium molybdate is added thereto. By this means a cobalt molybdate gel is precipitated on the alumina gel carrier. Catalyst preparations similar in nature to these and which can also be employed in this invention have been described in U. S. Patents 2,369,432 and 2,325,033.

Still other methods of catalyst preparation may be employed such as by impregnating dried carrier material, e. g. analumina-silica gel, with an ammoniacal solution of cobalt nitrate and ammonium molybdate. lPreparations of this type of cobalt molybdate catalyst are described in U. S. Patent 2,480,361. Inyet another method for preparing yimpregnated .imolybdate catalyst the carrier material may be rst impregnated with an aqueous solution of cobalt nitrate and thereafter impregnated with an ammoniacalmolybdate. Alternatively, the carrier may also be impregnated with both solutionsin reverse order. Following the impregnation of the carrier by any of the foregoing methods the material is drained, dried andnnally activated in substantially the same manner .as is employed `for the other catalysts. In the preparation of impregnated catalysts where separate solutions of cobalt and molybdenum are employed, it has been found. that it is preferable to .impregnate the carrier first with molybdenum, e. g., ammoniacal ammonium molybdate, and thereafter to impregnate with cobalt, e. g., aqueous cobalt nitrate, .rather than in reverse order.

In yet anotherV method for the preparation of suitable catalyst agel of cobalt molybdate can be prepared as describedhereinbefore for the unsupported catalyst, which gel after drying and grinding canbemixed Awith a ground alumina, alumina-silica orother suitable carrier together with a suitablepilling lubricantor binder which mixturecan then be pilled or otherwise formed into pills or largerparticles .and activated.

In yet another modification finely divided or ground molybdic oxidecanbe mixed with suitably ground carrier suchas.alumina, aluminasilica and the like in the presence of a suitable lubricant or bindery and thereafter pilled or otherwise .formed into larger agglomerated particles. These .pills or particles are then subjected to .a preliminary activation by heating to 600 C.,.for example, and. are thereafter impregnated with an aqueous .solution of' cobalt nitrate to deposit the cobalt thereon. .Afterdraining and drying the particlesare heated to about 600 C. to form the catalyst.

It is apparent from the foregoing description of the different types of cobalt molybdate'catalyst which maybe employed in this invention that we may-.employ either an :unsupported catalyst, in which case the active agents approximate 100% of the composition` or we mayemploy a supported catalyst wherein the active agents, cobalt and molybdenum oxides. will generally comprise from about 7 to 22% by weight of the catalyst composition. In all of the foregoing catalytic preparations it is desirable to maintain the molecular ratio of cobalt oxide as-CoO to molybdic oxide as M003 between about 0.4 and,5.0.

The desulfurization catalyst may be employed at temperatures between about 700 and 1000 F. However, it is preferable that temperatures in the range of about '750 to 850 F. be employed since .under theseconditions there is a minimum build-up of carbon on the catalyst and such conditions therefore permit long on-stream operating periods before regeneration becomes necessary, such as up to 200 hours 0r more. Suitable pressures range from about 500 to 10,000 and preferably from about 1000 to 2000 pounds per square inch. Feed rates from about 0.1 to 10.0 and preferably between 0.4 and 2.0 volumes of liquid feed stock per volume of catalyst per hour are employed in the desulfurization stage. About 1000 to 10,000 cu. ft. of recycle or fresh hydrogen are employedper barrel of feed.

The process of the present invention may be effected in any suitable equipment, a .particularly satisfactory system comprising a vfixed bed process in which the catalyst is disposed in oneor. more reactionzzonesv and .theunsaturated fraction to be hydrogenated is passed therethrough either in upwardor downward flow and. either concurrentlyl or countercurrently to a-stream of hydrogen. Fluidized type process in which the catalyst is carried into the `reaction zone bythe unsatu` rated fraction and/or hydrogen and maintained in a state of turbulence in the reaction lZone under hinderedsettling conditions,.moving bedtype process in which the `reactants are `passed concurrently or countercurrently toa moving bed of catalyst, or thefsuspensoid type operation in which the catalyst is .carried as a slurry in the reactants, may also be employed, Conventional processes using guard reactors to remove impurities from the charging stock prior tothe vhydrogenation treatment may be employed if desired.

The reforming catalyst of the present invention -is .a lplatinum catalyst preferably supported on alumina and which contains between about 0.5 and 8% by weight of combined halogen. Preferably the catalyst should contain between about 0.01. and 1.0% by. weight of platinum. Whilealumina is thepreferred support, it may also -beA supported on the carriers described hereinbefore in connection with cobalt molybdate catalysts. Iny thepreferred method for preparing the reforming .catalyst the halogen, which is preferably fiuorine or chlorine, is incorporated into the alumina .prior to comingling.,ordeposition of the platinum thereon. In the-preferred-modication alwetaalumina gel is'comingled withthe requisite amount of hydrogen chloride .and is thereafter comingledrin the. wet state with a solution of chloroplatinic acidwhich has been treated with hydrogenL sulfide. The composite yis filtered, dried and heated to a temperature between about 800 and 1200 Fpand preferably inthe presence of hydrogen.

The alumina `gel of this invention employed for `supporting theplatinum may be prepared by any ofthe methods well known 'inathe art, such as by treating any water-soluble aluminasalts with'eitheracids or bases in order to :cause the precipitation of aluminum hydroxide. In the preferred modification halogen-containing .aluminum salts. such as either aluminum .chloride or aluminum. fluoride are :treatedswith ammoniaV or otherI alkaline reagent to cause vprecipitation of thegel,y which .normally contains appreciable amountsfof halogenswhich are sufficiently .combined that they are removable onlyupon extended.

washing with fresh Water. Gels of this type may be .mixed `with .hydrogen sulded chloroplatinic acidmixture or,.alternatively, such .gels .may be treated at temperatures up to '1300 F. and thereafter impregnated with aqueous solutions of platinum-containing materials such as chloroplatinic acid.

Where the alumina gel is comingled with the platinum-containing compound it is ltered to a dry cake and dried at low temperatures such as in the range of 200 to 400 F. Following the initial drying the catalyst can be formed into pills, etc. in much the same manner as has been described hereinbefore in connection with the desulfurization catalyst.

The reforming step is carried out at temperatures between about 850 and 1050o F. and preferably in the range of 900 to 1000 F. Apressure in the range of about 50 to 1000 poundsper square inch is employed in conjunction with a liquid hourly space velocity between about 0.1 and and preferably between about 0.4 and 2.0. Recycle hydrogen amounting to between 1000 and 10,000 cu. ft. per barrel of charge stock is introduced with the charge stock to minimize carbon deposition. Operating cycles of up to several months may be employed between regenerations. i

Perhaps the process of this invention can best be understood by reference to the attached iigure.

The feed stock to the process may be any full range crude oil, reduced residuum, shale oil or the like. Such feed stock flows through line II to pump I2 thence through line I3, interchanger I4 and line 5 to preliminary fractionating tower I6. Fractionating tower I6 produces an overhead fractionated product to line I1 which is passed through lines |2| or |22 as will be described hereinafter. Topped feed stock flows from the bottom of fractionating tower I6 through line I8 to heater I9 wherein it is heated to a suitable coking temperature such as in the range of 800 to l000 F. Heated oil from heater I8 flows through line 20 to coking unit 2|.

Coking unit 2I is of any suitable conventional design for converting petroleum oils into gases and vapors boiling below a particular end point and producing a solid residue of coke. Such coking unit may be of the delayed coking type wherein coke is accumulated during a coking period of minutes to 2 hours, for and is then removed by suitable means with a stream of high velocity water in the well known hydraulic decoking method. In such cases coker 2| represents one of several coking vessels which are employed in sequence for-coking and decoking whereby a continuous ow of feed and products may be maintained. `In another modirlcation coker 2| may be of the continuous type wherein the oil is continuously passed through a bed of heated pellets of coke or other solid heat transfer material thereby depositing a layer of coke upon such pellets. The pellets containing added coke are removed and may be burned in order to recover the heat value of the coke and form additional hot pellets for further coking, or such pellets may be screened to recover the smaller pellets which are reheated and returned to the coking zone. In general such coking operations are effectedat relatively low pressures such as in the range of atmospheric to 100 pounds per square inch.

Coke is shown in the attached figure as being discharged from vessel 2| of the foregoing methods. Coker distillate is removed from vessel 2| through line 23 whence it passes through condenser 24 into line 25 and such as through line 22 by anyv thence to `fractionating tower 26. The overhead*- from fractionating-tower 26 passes through lines 21 and 28 into the fuel gas recovery system. Fractionating tower 26 discharges gasoline boiling range stock through line 29,*light gas oil through line S30 and heavy gas oil through line 3|. Light gas oil in line 30 may be joined with thegasoline in line 28 by opening valve 32. Light gas oil in line 30 may be joined with heavy gasoil in line 3I by opening valve 33.

In the preferred method of operation heavy gas oil from fractionator 26 passes through line 3| to heater 34 which is operated under suicient temperature and pressure to effect at` least a partial cracking of the heavy gas voil flowing therethrough. Cracked products from heater 34 flows through line 35 and are cooled in condenser 36 whence they flow through line 31 toy fractionating tower 38. Gaseous products of the cracking operation pass overhead from fractionating tower 38 through line '39 whence they join light gaseous cracked products from the coking step and flow through line 28 to the fuel gas recovery system. The heavy bottoms from fractionating tower 38 discharged through line 40, valve 4| and line lv42 to fuel oil storage not shown. In one modication of the invention the uncracked bottoms from fractionating tower 38 flow, upon the closing of valve 4|, through line 43, valve 44 and line 45 to pump 46 line 41 and join in line I8 the topped feed stock flowing to heater I8 as charge stock to coker 2|.

While a thermal cracker 34 is shown, such may be replaced with a catalytic cracker if such is desired. Catalytically cracked vapors are then supplied to condenser 30 and fractionating tower 38 as shown.

The gasoline boiling range stock from the cracking operation is withdrawn from fractionating tower 38 through line 48 and joins with the coker gasoline owing in line 29 whence the products ow through line 4S to heater 50. In the preferred modification, light gas oil' from fractionating tower 26 is caused to iiow from line 30, by the closing of valve 33 and the opening of valve 32, into line 29 vwhence it iiows through line 49 into heater 50. Recycle hydrogen is supplied with the charge to heater 50 through line 5| as described hereinafter. Ef-

fluent from heater 50 discharges through line 52A into desulfurization reactor 53 wherein it is contacted with a cobalt molybdate desulfurization catalyst of the type described hereinbefore. Desulfurization reactor eilluent iiows through line 54, condenser 55 and line tower v51.

The bottoms of fractionating tower 51 comprises desulfurized and hydrogenated light gas oil, the latter being suitable for diesel engine fuel without further refining. 51 is suitably controlled so ing point of the bottoms product is in the range of 400 to500 F., preferably. Bottoms product from fractionating tower 58, valve 59 and line 60 not shown.

Alternatively, the bottoms product from fractionating tower 58 may be recycled, if desired, to the cracking step. This is accomplished by Fractionating tower that the initial boil'- to diesel fuel storage closing valve 59 whence the product fiows from` line 58 through line 6I, valve 62 and line 63 into pump 64 whence it flows through line 85 and joins the charge stock to thermal cracking heater 34.

The overhead from fractionating tower 51 is whence they are forced into 56 into fractionating 51 flows through lineges-tree 1erefrilily controlled to ineludepr'piiegf pophe and lighter and is removed thf' ""lieftt'andvalief81 and line 8B whence-it "flows Vto the hy# drogen recovery system 'described'' hereinafter: The'partial closing of-valie 61 divertsla part 'of the gaseous stream froi'n line 68 through h'ydrogen analyzer Controller recorder' at which is in tlirn employedto ctate'rhotor valve 18 's described hereinafter.

The C4 to 40G-'500 F. fractioniiroiri'fratiate illg tower 51 floWS through line 1l to -heter 12i Recycle hydrogen from line' T3' joins thel gasoline range stock o'w'ihg to heater 12. Effluent of heater 'l2 flows through line 14 to catalytic reformer 15.

Catalytic reformer 15 employs a platinumalumiha reforming `catalyst of the type and une derl the conditions described' hereinbefor'e. Reformer efliuent hows through' line 1B, through heat exchanger 11 and line 18 int-o gas-liquid Separator 19. Hydrogen-rich gas from gas-liquid separator 19 flows through line 84 to the hy-` drogen recovery system described hereinafter. Liquid from gas-liquid separator 18 is removed under control of a liquid level control not shown' through line 80, heat exchanger 8|'a`n`d line 82 whence it vpasses into fractionating tower 83.l

, Fractionati-ng tower 83 is a rerun tower for the reformed gasoline boiling range product. Finished gasoline flows from fractionating tower 83-as the overhead through line 85 to product storage not shown. Bottoms'product' 'from fractionating tower 83 comprises small amounts of polymer and other miscellaneous products. Poly; mer' is removed from rerun tower 83 through line 8E; valve 81 and line 88 to polymer storage. If desired, such polymer may be recycled to the cracking step by closing valvev 81 Whereuponthe product flows .through line 89,v open valve 9|); line 8| when it entersy pump 92 and is forced into line 93 whence it flows through line `S'Sf'to heater charge flowing to heater 34.

Referring now more partieu'iarly to the hy"' drogen recovery system, hydrogenrich' gas from the desulfurization step in line 88 and hy'droger'reA rich gas from the reforming stepl line 84' arejoined in line 944 whence they flow through valve 85 and line 98 as charge stock to -hypersorp'e tion unit 98. The pinching of valve 85 causes" ay part of hydrogen-rich gas `in line 94 to' lhe' A diverted through hydrogen analyzer controller recorder 88 which may be employed to' aotuate motor valve 18 as described hereinafter:

Hypersorption unit 98" is ofthe type such as is shown in U. S. Patent 2,519,873 and consists of an adsorption zone, a rectification z'one,- astripper, charcoal rate controller anda l'ift line 9S.A Hydrogen-rich gasf'entering through line '91 passes upwardly through cold charcoal which absorbs Cr and heavier irachena-permitting the hydrogen to remain substantially unabsorbed; The hydrogen passes upwardly through' the-bedI and is removed through line |08'. The charcoal; containing adsorbed methane and heavier cone" stituents flows downwardly from the absorption' zone and enters the rectification zone, the lower portion of which comprisesl a' stripping zone.l The application of heat to the charcoal in4 the' stripping zone causes the gases to be' partially' desorbed from the-absorbentV and to passupwa'rde ly through the descending charcoal whence-it contacts cooler charcoal and causes: any exchange' of heavier for lighter constituentsin-the manner of rectification. Following rectification thef charcoal containing methane and heavier' cohstitient's'is thieafterheated to a temperature sufliiehttocau'sethe remaining gases to be lCle'- sorbetl whereupon Vthey 'are removed through line ll and pass to fuel gas storage rnot shown. Heated charcoal from the stripper hows through afchar'co'al 'regilator which controls the volume of 'charcoal di" 'ping therethrough per unit time to' sne'desird rate. Charooal'rom the charcoal regulator' 'passes downwardly through line |82 whence 'it is lifted through lift line 99 and is returned `to the to'p of the hypersorber. Such lifting may be effected i'n 'any suitable manner such as by gaslifting' or mechanical elevators. At the top of the'hyper'srfbr the heated charcoal is cooled to condition it for absorption of additidnal'gas.

another modification of the invention the hypersorber may be .employed in a side cutting operation so asl'to produce a C1-'C2-rich strea'ni through lr'ie |83 by opening valve |84 whence it `flows into line |85 to storage'. By employing this method a Cs-Hzsuich stream may be produced through line" iUi. This method is optional; hou/'even The gas iiowing in line 84 from the reformer will generally contain 8DL90 volume per cent of hydrogen and the gas flowing in line 68 from the desulfurizing unit will generally contain only 20-'50 volume per cent of hydrogen. Such gases may therefore be introduced into the hypersorb'er at two separate feed engaging zones. Thus by closing valve |4| and opening valve |42 about 20-"-5'0 volume per centv hydrogen is introduced in lower feed line 91 while 80-90 volume per cent hydrogen flows through valve |42 and upper feed line AHitt into 'an upper feed engaging zone. By this method leanl hydrogen is introduced lower in the tower so as to Contact a greater amount of charcoal and rich hydrogen is in`` tro'duced higher up so as to contact less charcoal; A- lower total flow of charcoal is needed to make a given separation thereby.

Hydrogen in line |88 flows through flow control-ler recorder |88 'which may be employed to actuate motor v'alve 18. Hydrogen thereafter flows through -line' |40 and line 13 to provide a recycle for the catalytic reformer and throughv line |48 andv line 5l to provide' a recycle for the.Y desulfuri'zationreactor.

The hydrogen content of the make gas froml reformer reactor 15 owing in line 84 is usually to 90' volume per ce'nt4 `and such gas may be employed as recycle gas to either the desulfurizationreactor 53,- orto reformer reactor 15 or to both. For this purpose valve |39 is openedI and' compressor |38 then takes suction on line' 84 and forcesV make gasv into line" |45) whence it passes to the reformer heater 12 or through line 5| to the desulfurizat-io'n reactor heater 5|).V

I-n another modificationof the invention the make gas'v from the desulfurizationV reactor 53 and the reformer reactor 15 are recycled, the former Abeing passedfrorn line 84 through valve |4 land-valve |138 to compressor |38l and the latter from line 84, valve |39 and compressor |38; For this purpose the net make gas is taken through valve 95 to'the hypersorher and processedto recover the hydrogen content and the hydrogen isreturnedA through line |88 and ori'- fice plate- I86 to` line |402` Referring now more particularly to fuel gas recovery' system, cracked; gases' from the ecker and thermal cracker new mrmighy une 2'8"@111' contain moderate amounts" of hydrogen! which `still another alternative, motor valve may be recovered if desired. Other waste refinery gases containing hydrogen, e. g., catalytic cracked oif gases, may be passed through line I09a by opening valve IIBa into line I I I which joins line |01 passing to compressor |08. Efuent gases from compressor |08 flow through line I 09 to gas-liquid separator IIO. Liquid from gas-liquid separator I I is removed through line II2 whence it ows to line 49 to join the heater charge to heater 50. Alternatively, compressed liquids may be discharged through line I 50 by opening valve I 5I. Gas from gas-liquid separator IIO passes overhead through line II3 and line IM whence they are sent to fuel gas storage not shown. If it is desired to recover hydrogen from such gas, it is then passed from line II3 into line II5, opened valve II-B and line II1 into line 91 whence it joins the charge to hypersorber 98 and is handled as described hereinbefore.

Referring now more particularly to the addition of straight run to the charge stock nowing to the catalytic reformer 15, straight run gasoline boiling range stock is introduced -through line |18 whence it ows through motor valve and line I I9 into line 1I. Where it is desired to augment the straight run stock with the straight run from the crude oil feed valve |20 may be opened whereby line I I8 is connected through line I2I with line I1 containing the topped straight run distillate of the crude.

The ilow of straight run to the catalytic reformer through line II8 is controlled by motor valve 10. Motor valve 10 in turn may be controlled through any one of three controllers. Hydrogen analyzer recorder controller 69 determines the hydrogen content of the gaseous components from the desulfurization reactor 53. Such gases are generally in the range of 25 to 50% by volume of hydrogen. Thus by setting the hydrogen controller recorder 69 to control the hydrogen concentration of the gas in line 86 at some definite Value, normally closed motor valve 10 will be opened to permit a greater flow of straight run to reformer as the hydrogen concentration of the off-gas of desulfurizer 53 decreases, and vice versa. Alternatively, motor valve 10 may be controlled by hydrogen analyzer controller recorder 98 which determines the hydrogen concentration of the combined off-gas from the desulfurizer 53 and the reformer 15 which flows through line 94. As 10 may be controlled by flow recorder controller IIlI which is actuated by the volume of hydrogen produced from hypersorber 98 through line I00. In all such cases or in other possible combinations the concentration of hydrogen or the volurne of hydrogen or both in the entire system or parts thereof may -be employed to control the amount of straight run supplied to the remainder to maintain the system in hydrogen balance. If, in the fractionation of crude oil in fractionater I 6, the amount of straight-run which is taken overhead and passed to reformer 15 via lines I1 and |22, heater 50, desulfurizer 53, fractionater 51, line 1I and heater 12 is sufficient to maintain the system in hydrogen balance, the hydrogen controls just described will cause motor valve 10 to remain closed; otherwise valve 10 will be opened as described above to permit further addition of straight-run gasoline through line II8.

In an alternative operation where it is desired to feed straight run from the topped crude to the desulfurization reactor, valve |20 may be closed so that such overhead distillate from fractionating tower I6 will then now through line I22 and opened valve |23 whence it flows through line |24 to line 49 and joins the heater charge to heater 50. Under this type of operation fractionating tower I6 may be operated to produce straight run gas oils overhead whereby such straight run stocks by-pass coking and cracking operations.

In still another modification the entire coker distillate in line 23 may be withdrawn through line |25, opened valve I 26 and line I21 whence it joins line 124 and flows as the charge stock in whole or in part to heater 50.

Perhaps the application of this invention can best be understood by reference to the following specific example.

Example end point 3090 Light gas oil, 400-700" F 5300 Heavy gas oil 3368 The heavy gas oil from the foregoing operation is subjected to thermal cracking whereupon there is produced the following products:

B./D. Pressure distillate, 400 F. end point 896 Fuel oil 2180 The pressure distillate from the cracking is combined with the pressure distillate and straight run and light gas oil from the coking step, making a total volume rate of 9286 B./D. which was then subjected to desulfurization at about 800 F. and 1500 p. s. i. with the addition of 3000 cu. ft. per barrel of hydrogen recycle. About 7,000,000 cu. ft. of hydrogen is consumed. The feed stock to the desulfurizer contains about 3.5% sulfur. A liquid feed rate of about 1.0 volume per volume of catalyst per hour is employed.

The catalyst is a cobalt molybdate type catalyst consisting essentially of 2% by weight of CoO and 10% by weight of M003 supported on 5% SiOz-% A1203, the carrier in coprecipitated form. The dried SiOz-AlzOa cogel is impregnated rst with an ammoniacal ammonium molybdate solution, dried, heated to 1200 F. for 2 hours, impregnated secondly with an aqueous solution of cobalt nitrate, dried and heated to l200 F. for 2 hours.

The olf-gas from the desulfurizer contains about 31% hydrogen and about 900,000 cu. ft. per day are bled from the system to prevent build-up of inerts and are processed to recover about 280,000 cu. ft. of hydrogen per day. The fractionated desulfurized product yields 4300 B./D. of diesel fuel.

The 5000 B./D. of 400 F. end point desulfurized gasoline from the desulfurizer are combined with 5000 B./D. of straight run gasoline of 400 F. end point quality and supplied to the reformer operating at about 950 F. and at a pressure of 750 p. s. i. Recycle hydrogen in the amount of about 4000 cu. ft. per barrel is supplied. A feed rate of about 2.5 Volumes of charge per volume of catalyst per hour is employed in the reformer.

The reforming catalyst is prepared by sulding amenace a chloroplatinic acid solution and mixing the resulting solution with an aqueous mixture of aluminum fluoride and ammonia, The precipitate is ltered and washed with dilute ammonia and dried. After heating to l200 F., the catalyst contains about 0.4% by weight of platinum and 4% by weight of halogens.

The reformer feed stock contains about 0.5% sulfur and the product .contains 0.001% sulfur. The product from the reformer amounts to about 9260 B./D. The net make gas from the reformer contains about 86.2% hydrogen and amounts to 7,780,000 cu. ft. per day. The make gas is processed to recover the hydrogen which is supplied to the desulfurization reactor.

I Orf-gases from the coker and cracker contain 37 tons per day of hydrogen suliide which is removed by scrubbing with an alkaline reagent and about 44 tons of hydrogen sulfide per day are recovered from the ofi-gas from the dcsulfurization unit'. Hydrogen could be recovered from these gases also.

It is apparent from the foregoing example of my invention that the entire hydrogen requirements for the desulfurization is met by the processing of make gas from the reforming step by adding straight run gasoline to the reforming step which is simultaneously upgraded. The term "st'raight-run gasoline is Well understood in the art 'as meaning `any gasoline fraction derived from crude oils by simple distillation, as distinguished from hydrocarbon cracking or reforming operations, Whether or not such straight-run gasoline fraction may have been subjected to conventional refining treatments such as acid treating, doctor sweetening, or other desulfurizing treatments,

Attempts to desulfurize theentire crude without inclusion of the coking step lead to excessive hydrogen consumption and prohibitive quantities of straight run gasoline are then required. In the present invention carbon-rich material is removed during the coking step With a consequent increase in the hydrogen to carbon ratio of the coker distillate. The amounts of hydrogen required to desulfurize and hydrogenate coker distillate is much lower than that required to process the crude itself. Furthermore, a considerable amount of thermal desulfurization takes place during coking to form hydrogen suliiide thereby decreasing the desulfurization load.

The liquid yield obtained in the three main steps is high and is usually about 80-90 volume per cent for coking, 9D-105 volume per cent for desulfuri- 'z'ation and 80-95 volume per cent for reforming. The overall yield of marketable liquid products based upon the crude is usually in the range of about 80 volume per cent. In addition a substanytial yield of coke and fuel gas is also obtained.

While hypersorption is the preferred method for effecting hydrogen recovery, other separation methods such as low temperature distillation, oil

absorption and the like may be employed.

In a hypersorption unit the charcoal must absorb the heavier constituent and transport them downwardly through the adsorption and rectification Zones. Hence the greater the concentration of the heavier constituent the greater the charcoal requirement. In the present invention the hydrogen-rich gas produced in the reformer is usually k80--90 volume per cent hydrogen and only a small amount of inerts need be removed therefrom. The olf-gas from the desulfurizer lusually contains between about -50 volume per c'en't hydrogen. It has been found that the handling of both hydrogen streams in a single hypersorber offers certain advantages in that the heavier constituents are diluted and the amount of rectification is decreased While neither the absorption nor the charcoal recirculation rate need be increased for a given purity of hydrogen production.

Cobalt molybdate is the preferred catalyst for the desulfurization reaction since it has been found that it gives a maximum upgrading of the stock for a minimum hydrogen consumption when it is used under the preferred operating conditions. Other catalysts are less eflicient for sulfur removal and especially for nitrogen removal, both of which elements are harmful in that they poison the reforming and especially the platinum catalysts. Olei'lns are mostly hydrogenated which minimizes carbon formation during reforming. Furthermore, other catalysts consume larger amounts of hydrogen in destructive hydrogenation to form methane, ethane and the like from which the hydrogen is not readily regenerable.

Platinum is the preferred catalyst for the reforming since it is very selective in its dehydrogenation and gives high hydrogen yields of high purity. Other catalysts give less hydrogen. While platinum is generally poisoned by processing even hydrogenated stocks, it has been found that when employed with cobalt molybdate, the removal of oxygen, nitrogen, sulfur, unsaturates, etc., is suiciently good that platinum catalysts can be operated satisfactorily.

However, where it is desired to employ other catalysts in this invention, it has been found that molybdenum oxides and suldes, tungsten oxides and suliides, chromium oxides, vanadium oxides and sulfides and combinations of `these elements with themselves or with such elements as copper, cobalt, iron, nickel and the like may be employed either supported or unsupported.

Where very large amounts of straight run gasoline are available for reforming, any of the aforementioned alternative desulfurization catalysts may be employed in the reforming step. Chromium-containing catalysts are particularly useful for this purpose.

The alternative catalysts for desulfurization and reforming are employed in substantially the same reaction conditions as that described hereinbefore for cobalt molybdate and platinum.

The foregoing disclosure of my invention is not to be considered as limiting since many variations may be made by those skilled in the art without departing from the scope or spirit of the following claims.

I claim:

l. In a two-stage process for refining hydrocarbons which includes passing said hydrocarbons in admixture With hydrogen into a catalytic hydrodesulfuriaation zone, withdrawing products from said hydrodesulfurization zone comprising hydrogen and hydrocarbons, separating a first hydrogen-rich gas from said products, recycling at least a part of said first hydrogen-rich gas to said hydrodesulfurization zone, separating a normally liquid desulfurized gasoline stock from said products, passing said gasoline stock in admixture with hydrogen into a catalytic hydroforming zone, withdrawing products from said hydroforming zone comprising hydrogen and reformed gasoline, separating a second hydrogen-rich gas from said products of said hydroforming zone, recycling a iirst part of said second hydrogen-rich gas to saidhydroforming zone, and passing a second recycle `stream of said hydrogen-rich gas to said hydroa platinum containing reforming catalyst is employed in said hydroforming zone.

3. A process according to claim 2 wherein said hydrodesulfurization zone is maintained at a temperature between about 750 F. and 850 F. and said hydroforming zone is maintained at a temperature between about 900 F. and 1000 F.

4. In a two-stage process for refining hydrocarbons which includes passing said hydrocarbons in admixture with hydrogen into a catalytic hydrodesulfurization zone, withdrawing products from said hydrodesulfurization zone comprising hydrogen and hydrocarbons, separating a first hydrogen-rich gas from said products, separating a normally liquid desulfurized gasoline stock from said products, passing said desulfurized gasoline stock in adlnixture with hydrogen into a catalytic hydroforming zone, withdrawing products from said hydroforming zone comprising hydrogen and reformed gasoline, and separating a second hydrogen-rich gas from said products of said hydroforming zone, the improvement which comprises maintaining an endogenous hydrogen balance within said process by including with said desulfurized gasoline stock supplied to said hydroforming zone a proportion of a straight run gasoline stock controlled to provide a net production of hydrogen in said hydroforming zone which is at least equal to the consumption of hydrogen in said hydrodesulfurization zone, recycling at least sufficient of said second hydrogen-rich gas to said hydrodesulfurization zone to balance the consumption of hydrogen therein, and recycling suiiicient of said first and second hydrogen-rich gases to said hydrodesulfurization and hydroforming zones to supply the total hydrogen requirements for each of said zones.

5. A process according to claim 4 wherein a Icobalt molybdate type desulfurization catalyst is employed in said hydrodesulfurization zone and a platinum-containing reforming catalyst is employed in said hydroforming zone.

6. A process for rening crude oil which comprises passing said crude oil into a coking zone maintained at a temperature between about 800 F. and 1000 F. wherein said crude oil is converted to coke, gases, and a Coker distillate having an increased ratio of hydrogen to carbon, separating said coker distillate, passing said coker distillate in admixture with hydrogen into a catalytic hydrodesulfurization Zone, withdrawing products from said hydrodesulfurization vzone comprising hydrogen and hydrocarbons,

separating' a 'rst hydrogen-rich gas from said products, recycling at least a part of said first hydrogen-rich gas to said hydrodesulfurization zone, separating a normally liquid desulfurized gasoline stock from said products, passing said desulfurized gasoline stock in admixture with hydrogen and a straight run gasoline into a catalytic hydroforming zone, withdrawing products from said hydroforming zone comprising hydrogen and reformed gasoline, separating a second hydrogen-rich gas from said products of said hydroforming zone, and recycling a part of said second hydrogen-rich gas to said hydroforming zone, the proportion of straight run gasoline admitted to said hydroforming zone being controlled so that there is a net production of hydrogen in said hydroforming zone which substantially balances the consumption of hydrogen in said hydrodesulfurization zone; and passing said net production of hydrogen to said hydrodesulfurization zone.

7. A process according to claim 6 wherein a cobalt molybdate type desulfurization catalyst is employed in said hydrodesulfurization zone and a platinum containing reforming catalyst is employed in said hydroforming zone.

8. A process according to claim 7 wherein said hydrodesulfurization zone is maintained at a temperature between about 750 F. and 850 F. and said hydroforming zone is maintained at a temperature between about 900 F. and 1000 F.

9. A process for refining crude oil which comprises passing said crude oil into a coking zone maintained at a temperature between about 800 F. and 1000 F. wherein said crude oil is converted to coke, gases, and a coker distillate having an increased ratio of hydrogen to carbon, separating a lower boiling fraction and a higher boiling fraction from said coker distillate, passing said higher boiling fraction through a cracking zone, withdrawing cracked products from said cracking zone, separating from said cracked products a lower boiling fraction and a higher boiling fraction, passing said lower boiling fraction from said coker distillate and said lower boiling fraction from said cracked products in admixture with hydrogen into a catalytic hydrodesulfurization Zone, withdrawing products from said hydrodesulfurization zone comprising hydrogen and hydrocarbons, separating a rst hydrogen-rich gas from said products, recycling at least a part of said iirst hydrogen-rich gas to said hydrodesulfurization zone, separating a normally liquid desulfurized gasoline stock from said products, passing said desulfurized gasoline stock in admixture with hydrogen and a straight run gasoline into a catalytic hydroforming zone, withdrawing products from said hydroforming zone comprising hydrogen and reformed gasoline, separating a second hydrogen-rich gas from said products of said hydroforming zone, and recycling a part of said second hydrogen-rich gas to said hydroforming zone, the proportion of straight run gasoline admitted to said hydroforming zone being controlled so that there is a net production of hydrogen in said hydroforming zone which substantially balances the consumption of hydrogen in said hydrodesulfurization zone; and passing said net production of hydrogen to said hydrodesulfurization zone.

l0. A process according to claim 9 wherein a cobalt molybdate type desulfurization catalyst is employed in said hydrodesulfurization zone and a platinum containing reforming catalyst is employed in said hydroforming zone.

11. A process according to claim 10 wherein said hydrodesulfurization zone is maintained at a temperature between about '750 F. and 850 F. and said hydroforming zone is maintained at a temperature between about 900 F. and 1000 F.

12. A process according to claim 11 wherein the liquid remaining after said separating of a normally liquid desulfurized gasoline stock from -said products oi said hydrodesulfurization zone is recycled to said cracking zone.

13. A process according to claim 11 wherein the gases produced during coking and during cracking are processed to recover hydrogen and at least a part of said hydrogen is passed into said hydrodesulfurization zone.

14. A process for rening crude oil which comprises passing said crude oil into a coking zone maintained at a temperature between about 800 F. and 1000 F., wherein said crude oil is converted to coke, gases, and a coker distillate having an increased ratio of hydrogen to carbon, separating a lower boiling fraction and a higher boiling fraction from said coker distillate, passing said higher boiling fraction through a cracking zone, withdrawing cracked products from said cracking zone, separating from said cracked products a lower boiling fraction and a higher boiling fraction, passing said lower boiling fraction from said coker distillate and said lower boiling fraction from said cracked products in admixture with hydrogen into a catalytic hydrodesulfurization zone, withdrawing produpts from said hydrodesulfurization zone comprising hydrogen and hydrocarbons, separating a rst hydrogen-rich gas from said products, separating a normally liquid desulfurized gasoline stock from said products, passing said desulfurized gasoline stock in admixture with hydrogen and a straight run gasoline into a catalytic hydroiorming zone, withdrawing products from said hydroforming zone comprising hydrogen and reformed gasoline, separating a second hydrogen-rich gas from said products of said hydroforming zone, passing at least a part of said first hydrogen-rich gas into an adsorption zone, adsorbing the more readily adsorbable constitutents from said hydrogen-rich gas in constituents from said hydrogen enriched hydrogen from said adsorption zone, recycling a first part of said enriched hydrogen to said hydrodesuliurization zone and passing a second part of said enriched hydrogen to said hydroiorrning zone, the proportion of straight run gasoline admitted to said hydroforming zone being controlled so that there is a net production of hydrogen in said hydroforming zone which substantially balances the consumption of hydrogen in said hydrodesulfurization zone; and passing said net production of hydrogen to said hydrodesulfurization zone.

l5. A process according to claim 14 wherein the gases produced during coking and during cracking are processed to recover hydrogen and at least a part of said hydrogen is passed into said hydrodesulfurization zone.

16. A process for refining crude oii which comprises passing said crude oil into a coking zone maintained at a temperaturebetween about 800 F. and 1000 F. wherein said crude oil is converted Vto coke, gases, and a coker distillate having an increased ratio of hydrogen to carbon, separating a lower boiling fraction and a higher boiling fraction from said coker distillate, passing said higher boiling fraction through a cracking zone, withdrawing cracked products from said cracking zone, separating from said cracked products a lower boiling fraction and a higher boiling fraction, passing said lower boiling iraction from said coker distillate and said lower boiling fraction from said cracked productsl in admixturewith hydrogen into a catalytic hydrodesuliurization zone, withdrawing products from said hydrodesulfurization zone comprising hydrogen and hydrocarbons, separating a first hydrogen-rich gas from said products, separating a normally liquid desulurized gasoline stock from said products, passing said desulfurized gasoline stock in admixture with hydrogen and a straight run gasoline into a catalytic hydroforming zone, withdrawing products from said hydroforming zone comprising hydrogen and reformed gasoline, separating a second hydrogenrich gas from said products of said hydroiorming zone, passing at least a part of said rst hydrogen-rich gas and said second hydrogen-rich gas into an adsorption zone, adsorbing the more readily adsorbable constituents from said hydrogen-rich gases in said adsorption zone, removing enriched hydrogen from said adsorption zone, recycling a rst part of said enriched hydrogen to said hydrodesulfurization zone and recycling a second part of saidenriched hydrogen to said hydroiorming zone, the proportion of straight run gasoline admitted to said hydroforming zone being controlled so that there is a net production of hydrogen in said hydroforming zone which substantially balances the consumption of hydrogen in said hydrodesulfurization zone; and passing said net production of hydrogen to said hydrodesulfurization zone.

17. A process according to claim 16 wherein the gases produced during coking and during cracking are processed to recover hydrogen and at least a part of said hydrogen is passed into said hydrodesulfurization zone.

References Cited in the le of this patent UNITED STATES PATENTS Number Name Date 2,273,224 Schulze Feb. 17, 1942 2,352,059 Woog June 20, 1944 2,417,308 Lee Mar. 11, 1947 2,479,110 Haensel Aug. 16, 1949 2,487,466 Nahin Nov. 8, 1949 2,498,559 Layng Feb. 21, 1950 2,500,146 Fleck et al Mar. 14, 1.950

Claims

1. IN A TWO-STAGE PROCESS FOR REFINING HYDROCARBONS WHICH INCLUDES PASSING SAID HYDROCARBONS IN ADMIXTURE WITH HYDROGEN INTO A CATALYTIC HYDRODESULFURIZATION ZONE, WITHDRAWING PRODUCTS FROM SAID HYDRODESULFURIZATION ZONE COMPRISING HYDROGEN AND HYDROCARBONS, SEPARATING A FIRST HYDROGEN-RICH GAS FROM SAID PRODUCTS, RECYCLING AT LEAST A PART OF SAID FIRST HYDROGEN-RICH GAS TO SAID HYDRODESULFURIZATION ZONE, SEPARATING A NORMALLY LIQUID DISULFURIZED GASOLINE STOCK FROM SAID PRODUCTS, PASSING SAID GASOLINE STOCK IN ADMIXTURE WITH HYDROGEN INTO A CATALYTIC HYDROFORMING ZONE, WITHDRAWING PRODUCTS FROM SAID HYDROFORMING ZONE COMPRISING HYDROGEN AND REFORMED GASOLINE, SEPARATING A SECOND HYDROGEN-RICH GAS FROM SAID PRODUCTS OF SAID HYDROFORMING ZONE, RECYCLING A FIRST PART OF SAID SECOND HYDROGEN-RICH GAS TO SAID HYDROFORMING ZONE, AND PASSING A SECOND RECYCLE STREAM OF SAID HYDROGEN-RICH GAS TO SAID HYDRODESULFURIZATION ZONE, THE IMPROVEMENT WHICH COMPRISES MAINTAINING AN ENDOGENOUS HYDROGEN BALANCE WITHIN SAID PROCESS BY MIXING WITH SAID DESULFURIZED GASOLINE STOCK SUPPLIED TO SAID HYDROFORMING ZONE A PROPORTION OF A STRAIGHT-RUN GASOLINE STOCK CONTROLLED TO PROVIDE A NET PRODUCTION OF HYDROGEN IN SAID HYDROFORMING ZONED SUBSTANTIALLY EQUAL TO THE CONSUMPTION OF HYDROGEN IN SAID HYDRODESULFURIZATION ZONE, AND PASSING SAID NET PRODUCTION OF HYDROGEN TO SAID HYDRODESULFURIZATION ZONE.