US2958643A

US2958643A - Two-stage catalytic conversion process for producing naphthalene and an aromatic gasoline from cycle oils

Info

Publication number: US2958643A
Application number: US606822A
Authority: US
Inventors: Bernard S Friedman
Original assignee: Sinclair Refining Co
Current assignee: Sinclair Refining Co
Priority date: 1956-08-29
Filing date: 1956-08-29
Publication date: 1960-11-01
Anticipated expiration: 1977-11-01

Description

GAS

Nov. 1, 1960 B. S. FRIEDMAN TWO-STAGE CATALYTIC CONVERSION PROCESS FOR PRODUCING NAPHTHALENE AND AN AROMATIC GASOLINE FROM CYCLE OILS Filed Aug. 29, 1956 1') LL! 2 2 2 9 5 1, 5 2% 5% 1 2 O r- I E 55 92 :1 25% O 2 E O 2 2 2 m 2 E SEPARATOR 8% REACTOR 3 g0 95% m g n x E .J I w 9 (O a g A SEPARATOR :5

d] IN REACTOR M Q r. Id 9 J0 O INVENTOR BERNARD S. FRIEDMAN BY W44 Mg United S at A TWO-STAGE CATALYTIC CONVERSION PROCESS FOR PRODUCING NAPHTHALENE AND AN AROMATIC GASOLINE FROM CYCLE OILS Bernard S. Friedman, Chicago, Ill., assignor to Sinclair Rfelfi/lnlng Company, New York, N.Y., a corporation ame Filed Aug. 29, 1956, Ser. No. 606,822

9 Claims. (Cl. 208-60) This invention relates to the production of commercial grade naphthalene and, more specifically, to the product1on of commercial grade naphthalene and high octane aromatic gasoline. The present invention is particularly concerned with a two-stage hydroconversion process in which a light cycle oil derived from petroleum is converted to naphthalene, alkylated naphthalenes and high octane number aromatic gasoline through the utilization of certain hydroforming catalysts in the presence of by drogen. This application is a continuation-in-part of my copending application Serial No. 434,059, filed June 2, 1954, now abandoned.

High octane gasoline and naphthalenes are highly desirable products which can be derived from petroleum hydrocarbons, and many refiners are in search of new methods of producing these materials from inexpensive charge stocks. The present invention is concerned with a two-stage process for manufacturing these products through catalytic conversion of light cycle oil stocks which are usually considered too refractory to be employed 'as feeds in conventional catalytic cracking procedures since the amount of coke formed on the catalysts at the conditions necessary for cracking is prohibitive.

I am aware that the production of naphthalene from cycle oils has been proposed, for instance, through utilization of the water gas reaction or hydropyrolysis. Although such processes may give desirable amounts of naphthalene, other valuable constituents of the charge stock are converted to less valuable products such as gas due to the severe conditions needed to effect the dealkylation of the alkyl naphthalenes in the cycle oils. In the present invention, I have devised a method which affords good yields of naphthalene from light cycle oils and at the same time transforms non-naphthalenic constituents to valuable high octane number aromatic gasoline.

Although my process will be defined below as a twostage hydroconversion of light cycle oil in the presence of non-carbon based hydroforming catalysts at certain reaction conditions, these conditions can be selected or varied within the ranges disclosed to emphasize particular chemical reactions. In the first stage (StageI) of the present process the reaction conditions can be maintained to convert the major portion of the non-naphthalenic constituents of the feed stock to; aromatic and non-aromatic gasoline with a minor amount or substantially none of these constituents being aromatized to the C to C range. The gasoline is then separated and the remaining aromatic-rich oil containing naphthalenic derivatives is charged to the second or more severe reaction (Stage II) of the process to produce good yields of commercial grade naphthalene, alkylated naphthalenes and additional amounts of aromatic gasoline.

Alternatively, the Stage I conditions can be selected to effect wholly or predominantly an aromatization re action with little or no aromatic and non-aromatic gasoline being produced. Although it is possible to produce no substantial amount of gasoline in the first stage which would obviate the desirability of a gasoline. removal 2,958,64 Patented Nov. 1, 1960 operation, this stage will in most instances, produce at least a small amount of gasoline which should be removed from the higher boiling aromatic concentrate before it is subjected to the second stage reaction. In this variation of the system among the products from the second stage are again naphthalene, alkylated naphthalenes and high octane aromatic gasoline. In commercial practice, an operator can vary the first stage reaction conditions within the prescribed limits as he chooses; however, in each instance the ultimate products are essentially the same. It may even be desirable to select the conditions for the first stage reaction so that a substantial amount of gasoline is made; yet non-naphthalenes in the feed stock could also be converted in substantial amounts to C to C aromatics.

In the initial stage the reaction is regulated in severity so that valuable non-naphthalenic constituents of the feed are converted to gasoline with a minimum loss through production of less valuable gas. In general, the reaction temperature can vary from about 900 to l200 F., while the pressure is maintained at least about atmospheric, for instance atmospheric to 1000 p.s.i.g. or higher, preferably less than 1000 p.s.i.g. The preferred reaction temperature is from about 900 to 1050 P. which affords especially high octane gasoline as a product. The space velocity, that is the volume of liquid hydrocarbon per hour per volume of catalyst (LHSV), can vary from about 0.1 to 20 LHSV; however, it must be correlated with temperature in order that neither too little nor too much conversion is effected in the first stage reaction zone. The severity of the first stage reaction is controlled so that the aromatic oil (product boiling above 400 F.) contains at least about 40 weight percent, preferably at least about 60 weight percent, of naphthalenic aromatics and not more than about 30 Weight percent, preferably not more than about 20' weight percent of non-aromatic components. Should insufficient naphthalenic aromatics be present in the aromatic oil derived by a single pass through the first stage reaction zone, this oil could be recycled to Stage I for further conversion. In this stage the severity is not increased materially beyond the point where more than about 10 weight percent of naphthalene is in the aromatic oil. Also it is preferred to operate the first stage reaction so that the yield of liquid product is over 70 weight percent of the feed. The space velocity will usually be from about 0.5 to 8 LHSV, preferably about 1 to 3 LHSV. A space velocity of about 2.0 LHSV is particularly preferred in producing high octane gasoline. The temperature and space velocity employed will be affected by the extent of conversion desired in the first stage which in turn is a function of the choice of catalyst, the hydrogen partial pressure and to some extent the total pressure and the degree of catalyst deactivation or the amount of coke deposited on the catalyst at a given time in the processing cycle.

Sufiicient hydrogen must be present in the first stage reaction zone to effect the reaction and maintain the activity of the catalyst through decreasing the amount of coke lay-down. In the first stage of the process the use of about 3 to 10 or 20 mols of hydrogen for each mol of feed stock produces satisfactory results. The use of about 5 to 6 mols of hydrogen per mol of feed is preferred, and the benefits derived by using more than about 10 mols of hydrogen usually do not justify the increased expense. of course, the specific amount of hydrogen employed will vary with the feed stock charged and the reaction conditions observed. Under certain select conditions of temperature and pressure, e.g. 1000 F. and 50 p.s.i.g., operating with recycled tail gas, the net consumption of hydrogen may be reduced so that an extraneous source of hydrogen is not required. Also,

silica-magnesia, or silica-zirconia.

gasoline and the aromatic-rich oil.

3 when operating with recycled tail gas, hydrogen sulfide is preferably removed from such gases.

The catalysts which can be used in both stages of the present process are non-carbon based hydroforrningcatalysts which can contain chromium, molybdenum, tungsten, cobalt, or vanadium (or mixtures of these) preferably deposited as oxides on non-combustible carriers such as alumina, titania, thoria, zirconia, silica, silica alumina, Platinum, palladium, rhodium, and other rare metals may be employed preferably'in the metallic state supported on the same noncombustible type of carriers. In operating my process on .a cyclic or continuous basis, the catalyst can be regenerated at intervals by treatment with air or oxygen at temperatures above 900 F.

The preferred catalyst is chromia-alumina since it affords an exceptionally selective action leading to the formation in good yields of thehigh octane aromatic Molybdena-alumina catalysts are, for example, not as desirable as chromiaalumina, since the former may produce more carbon and sulfur on the catalyst in both Stage I and Stage II ,of the process. chromia-alumina catalysts which are generally known in the petroleum processing field can be .utilized; and these catalysts usually contain from about 1.0% to about 25% or more by weight of chromia. The activity of the chromia-alumina system can be appreciably enhanced by the use of promoters which can in clude for instance silica, berryllium, boron, potassium and cerium. A specific catalyst found effective contains .about 12% chromia, 86% alumina, and about 2% magnesium oxidejas a promoter.

The feed stocks which can be employed in the initial reaction are petroleum light cycle oils boiling generally in the range from about 400 to 650. F. Also, aromatic extracts of these stocks can be used. The light cycle oils usually are composed of at least about 40 weight percent of non-aromatics and these may comprise about 50 to 65% or more of the oil. The feed stock can be --desulfurized but this is not considered essential as a -desulfurization reaction is effected in each stage. of the .present process and most of the sulfur present in the feed 'is removed through conversion to hydrogen sulfide in the case of chromia catalysts, or in part deposited on the catalyst as metallic sulfide, in the case of molybdenum catalysts.

In the second stage of the present process the aromatic oil from the initial reaction is treated under more severe conditions to insure the production of (a) sufiicient yield of naphthalene in highconcentration in its boiling range, -(b) conversion of benzenoid aromatics and non-aromatic .hydrocarbons to highly aromatic gasoline, and (c) to elfect removal of nitrogenandsulfur-containing impuri- .ties from the desired products. The temperature in the .secoud stage reaction can vary from about 900 to 1200 F. whilethe space velocity will generally be from about 0.1 to 3 LHSV. The reaction pressure and amount of :hydrogen supplied can be of the same range as specified for the initial stage of the process providing that one .or more of the operating conditions is considerably more severe than employed in Stage I. Compared with Stage I the desired severity of the-Stage II reaction may be effected by raising the temperature by at least about 125 obtained by a partial change of more than one of these variables in the indicated directions. It is preferred to operate the second stage ofmy process so that the liquid product yield is over about 70 weight percent of the total fee d to this stage and the product fraction boiling between 400 to 460 F. contains at least about 40 weight percent naphthalene and at least about 50 weight percent total aromatics.

The products of the second stage reaction include aromatic gasoline, naphthalene, beta-methylnaphthalene and other alkyl-naphthalenes. These products can be separated by fractional distillation. The naphthalene and beta-methyl-naphthalene are of high purity, substantially free from sulfurand nitrogen-containing impurities, and are practically completely aromatic. Beta-methylnaphthalene and the other alkylnaphthalenes can be converted to naphthalene as for instance by recycling to the second stage of my process the aromatic oil boiling above about 400 F. from which the naphthalene has been separated. 2,6-dimethylnaphthalene may be isolated by cooling, filtering, and recrystallization from the 'dimethylnaphthalene fraction either from the first or second stage products. Also, when betamethylnaphthalene and/or 2,6- dirnethylnaphthalene .aswell as naphthalene are separated from thereaction product of the second stage, the remaining oil boiling'above about 400 F. can be re 'cycledto the second stage reaction.

When utilizing a chromia-alumina catalyst in the re- .action of Stage I'of this invention, I prefer the severity in operating conditions in ,the first stage to be equivalent .toor defined by 1000 to 1100 F., -2 LHSV and 200 to 500 p.s.i.g. hydrogen partial pressure. In the second stage when employing thisccatalyst the preferred severity is equivalent to .or. ;definedxby 1050 to ll50 F.,

0.5 LHSV and 200to 5.00 p.s.i,.g. hydrogen partial. pres- :sure. Different catalysts do, .of course, re'quire differ- -.ent operating conditions;;to produce the best results.

Thus, if a molybdena-alumina catalystzbe employed the preferred severitiesinthe separate stages are, respectively, equivalent to. ,or definedzby-a900 to 1000 F., 2

LHSV, 50 to 400 p.s.i.g.hydrogen partial pressure and 900 to 1000 F., 1.0'LHSV,; 200 to'500 p.s.i.g. hydrogen partial pressure.

Thev process of the present invention can be effected,

for instance, in asingle reactor system or in a system including'tworeactors. Intthe single reactor system the initial stage will be effected and. the products separated as by fractionation. Thearomatic-rich oil is collected and then converted in a blocked-out operation to naphthalene and additional amounts of aromatic gasoline in the reactor after ithe'desiredamount of light cycle oil has been processed. In a two-reactor system the separate stages of the process can be conducted in the separate reactors. Also, in any reaction system provision can be made for regenerating the catalyst. 'Also, both stages may be conducted with fluidized or moving bed type of catalytic reactors.

It is not possible to obtain satisfactory yields of gasolineand naphthalene by operating with one reactor continually recyclingthe 400 F. plus product from which naphthalenehas been recovered. In the first -place, operatingconditions severe enough to effect demethylation off-methylated naphthalenes are much too severe for converting the non-aromatics to gasoline, and will produce more dry 'gas than the economics will per- .mit. Conversely .when'the conditions are optimum for gasoline production, they are not severe enough to deconcurrently with the cycle oil directly to the initial stage 'reactorto produce the needed hydrogen while being reformed.

()ne method of operation in the present invention consists in using the catalyst successively for reforming and the first and second stage reactions. Thus naphtha is passed over the freshly regenerated catalyst under hydroforming conditions, e.g., 0.5 LHSV, 950 F., 200 p.s.i.g. H recycling 4 to 6 mols of H At the end of one-half to two hours, when the reforming activity has diminished because of mild coke lay-down, the naphtha feed is cut out and replaced with cycle stock (first stage). The temperature may be raised to effect the desired reactions. Finally, after one to three hours on the first stage the reactor is switched to the second stage by feedingaromatic oil from the first stage. At the end of this operation the catalyst is regenerated and readied for the hydroforming step.

In the initial stage of my process as the operating temperature is raised, the dry gas, carbon make and con- Ma rate of about 2 liquid volumes per hour per volume of catalyst: API/60" F. 21.8 Percent S Wt. percent-.. 2.38

5 Olefins wt. percent 26.8 Aromatics wt. percent.. 51.2 Br No. 20.9 N 1.5251 B.P. Initial F 422 50% F 490 End-point F 568 At the end of four one-hour runs, the products were separated and analyzed on an average basis as follows:

version to gasoline increase. Also the yield of recycle Product Wt men; oil decreases but its aromaticity increases. As shown by o't i eed the following tabulation the initial stage aifords advantages over the thermal cracking of the light cycle stocks Liquid 79.4 conventionally effected. The gasoline produced in the gg ifi f-g reaction is higher in octane number (e.g. 86 to 95 vs. 73) galrfbon on o t tuy st .7 and the by-product liquid is more valuable because of u a its low boiling range and viscosity and its utilization as 5 100-6 a feed stock for naphthalene production.

Table I Feed (Light Cycle Oil):

API gravity, 60 F 21. 6 21.8 21.8 Percent S 1. 96 2. 38 2. 38 47.9 51.2 51.2 444 422 422 538 490 490 637 568 568 899 1,000 1,050 560 400 409 2 2 H; mol ratio.. 1 5/1 Catalyst Thermal CrzOa/AlzOa CraOa/AlzO: Products (Wt. percent):

Dry gas 6. 4 6. 6 13. 6 aka 0. 6 1. 3 Gasoline (oi-400 F.) 29.1 29. a 33.1 Research method octane no. neat (approx. 5 lbs. Reid 73 86 95 Aromatic oil (400-600 F.) 57.7 45.6 API gravity at 60 F (17.8) (13. 7) Tar 64.5 3.5 2.9 API gravity at 60 F (4. l)

One form of the present invention can be described with reference to the drawing which is a diagrammatic, simplified flow sheet of a system comprising separate reactors for the first and second stages, respectively, of

my process.

In the drawing, light cycle oil feed enters by way of line 1 into reactor 2, which contains the hydroforming type catalyst. A hydrogen-rich gas can enter reactor 2 by way of line 3 and the reaction product is conveyed in line 4 to separator 5. Gas is withdrawn from separator 5 through line 6. Aromatic gasoline is taken by way of line 7 and the bottoms are removed in line 8. An aromatic-rich oil is drawn from separator 5 by way of line 9 and passed to the second stage reactor 10 containing hydroforming catalyst. A hydrogen-rich gas enters reactor 10 by way of line 11 and the reaction product passes through line 12 to separator 13. Products withdrawn from separator 13 include aromatic gasoline, naphthalene, methyl naphthalene, other naphthalenics, and bottoms.

The following specific examples will serve to illustrate the present invention but they are not to be considered limiting.

EXAMPLE I A regenerated chromia-alumina catalyst (12% Cr O 2% MgO) was charged to a reactor and brought to a temperature of about 1050 F. A catalytic light cycle oil of the following analysis was charged to the reactor The liquid product was subjected to fractional distillation and produced the following:

Product:

Gasoline (C to 400 F.) 33.4 wt. percent of Feed (0.037; sulfur, 58.1% aromatics, 9.7% olefins).

Aromatic oil (400-600 F.):

Wt. percent of feed PIT/60 F Wt. percent olefin 13.6. Wt. percent aromatics. 84.9.

1.575. Specific dispersion 269. Bottoms:

Wt. percent of feed 1.8.

1 Research method octane number of O to 400 F. gasoline was 95.8.

This yield includes 0 from gas.

An aromatic-rich oil stock produced according to the specific procedure noted above but analyzing:

Boiling point F 400-600 API/60 F. 12.7 Specific dispersion 270 N 1.5740 Wt. percent aromatics 81.3 Wt. percent olefins 13.2

was charged 'at a space velocity of 0.55-volumes per hour per volume of catalyst to a reactor containing a chr'ornia-alurnina catalyst (12% Cr O 2% MgO). The catalyst was maintained at a temperature of about 1l09 The' liquid product was subjected to fractional distillation and produced the following:

Product: and a p e of 400 P- -s and hydrogen Was P Gasoline (C4 e R) H Wt, percent of Feed. plied to the reactor at a ratio of 5.7 mols per mol of v (64.1% aromatics, 2.4% feed. The reaction continued for 2 hours and the fol- Aromatic on 400 600F) 01mm)- lowing products were obtained, Wt. percent of Feed 4e. ARI/60 F 16. Wt. percent olefin 5.3. Wt. Percent percent aromat 83.7. Product of Aromatic i l-5620 011 Feed Specific dispersion 258.

Bottoms:

Wt. percent of Feed 2.1. 8 1 Research method octane number of C t 5 o 400 F. gasoline was 93.8. gag 6 This yield includes 0 from gas. Dry gas 19. 3 Carbon on catalyst 4. 34 Sulfur on catalyst.

TotaL. 102-04 Wt. Percent Wt. Percent of Feed of 400-600 F. The liquid product was distilled and the following prod- Pmduc ucts were separated: Ngnh "1mm 3- 05 6'1 Alkylnaphthalenes 23. 2 47. 3 Wt. Percent Product of Aromatic Oil Feed 0 The aromatic-rich oil stock boiling 400600 F. was gg 5 1 292 gig-g charged at a space velocity of about 1.0 volume per hour fl-methylnap 16.6 per volume of catalyst to a reactor containing a molyb- 3335 naphthalem 3g dena-alurnina catalyst 10.14% Moo -5% SiO The catalyst was maintained at about 985 F. and a pressure Research method octane number of C to 400 F. gasoline was equivaof of hydrogen and hydrogen was Supphed lent to iso;ootane +0.54 ml Eetraethyl lead. to the reactor at a ratio of 5.9 mols per mol of feed. ggl f g gjggf igfigf The reaction was continued for one hour and the follow- The results of the two-stage process are summarized mg Products were Obtamed: below:

Stage I Stage II Combined Wt Percent Yields 1 Product of Aromatic Oil Feed Run No 14 20 40 Temp. F. (400 lbs. p.s.i.g. H2)... 1,050 Liquid 73,4 LHSV, approx 2 0 in g 1. 8 Prtzdufts (Wt. percent feed to Dry g 14.0 S 91) I Carb on o tal st 10.9 Gasoline, C4 to 400 F. (V01. sulfu i on ca t aly s tn 0.2

percent) 33.40107) 21.0 46.9 (54. 3) Percent Arorn. (C to 400) 58.1 100 1m 3 Percent S (C5 to 400) 0.03 R.M.O.N'. (C5 to 400) 95.8 106.5 98.7 M.M.0.N. (0 to 400). 84. 106.5 90. 4 Heavier liquids fig% fi f g f:::: ff 2f "iii The liquid product was distilled and the following prod- Methylnaphthalene. 16. 6 ucts were separated: Heavier oil 1. 8 6. 3 5. 4 Carbon on catalyst. 1. 7 4. 3 4. 4 Dry gas 13. 6 19. 3 25. 6

Ultimate yields based on original light cycle oil feed and recycling F a alkylnaphthalenes to extinction in Stage II to produce maximum amount Product 5 of naphthalene. All percentages noted above are calculated on a weight 1 69 basis unless otherwise specified.

EXAMPLE II aS(%lt]111e1(C4t0400F.) 19.5 a aene 22.7 A regenerated molybdena-alumina catalyst (10-1 aniistli lrra hthelene 15.7 M00 on Al O -5% SiO was charged to a reactor and gamma naphthalemc Pmductwa ottoms. 5.8 brought to a temperature of about 950 F. The same feed employed in Example I was charged to the reactor at a rate of about 2 liquid volumes per hour per volume of catalyst. Hydrogen was supplied to the reactor at the rate of 4.92 mols per mole of feed, under a pressure EXAMPLE III of 400 p.s.i.g. At the end of 3 l-hour runs, the products were Separated and analyzed on an average basis as 'As an example of severity control in Stages I and II follows: of my process a number of operations were conducted using the catalyst of Example I at a variety of tem- Pmduct wtpercent peratures and space velocities. The light cycle oil feed OfFeed approximated that of Example I and the hydrogen was supplied to the reactor at the rate of 7 mols per mol of feed and the total pressure was 700 p.s.i.g. In these tests it was determined that at the reaction conditions 4.7 'v I sulfur on catalyst Lo gi en the following temperatures and space velocities could be employed; however, when using these conditions they should be selected so that the Stage II reaction 7 ls'more severe than that of Stage I" as noted above.

Stage I, Ap- Stage II, Approximate proximate Temperature, T. Range of Space Maximum Velocity, Space Veloc- LHSV ity, LHSV 0. 2 to 0. 5 0. 1 0.25 to l. 0. 2

0. to 2. 0 0. 0. 9 to 4. 5 0. 7 1. 7 to 9. 5 1.2 3 to 20 2. 5

1 The minimum space velocity in Stage II could be about 0.1 LHSV at each temperature as a slower rate of feed would not be commercially easible.

I claim: 1

1. The method of producing naphthalene and high octane aromatic gasoline which comprises contacting in a first stage in the presence of hydrogen a petroleum light cycle oil consisting essentially of aromatics and about 40 to 65 percent of non-aromatic components with a hydroforming catalyst having a non-combustible base at a temperature of about 900 to 1200 F., a space velocity of about 0.1 to 20 LHSV and a pressure of at least about atmospheric while efiecting conversion to an aromatic oil boiling above 400 F. and containing at least about 40 weight percent of naphthalenic aromatics and not more than about 30 weight percent of non-aromatics, said aromatic oil containing not more than about 10 weight percent of naphthalene, contacting in a second stage under more severe reaction conditions the aromatic oil with a hydroforming catalyst having a non-combustible base, in the presence of hydrogen and at a temperature of about 900 to 1200 F., a space velocity of about 0.1 to 3 LHSV and a pressure of at least about atmospheric to form naphthalene and aromatic gasoline, and separating the gasoline and naphthalene.

2. The method of claim 1 in which the aromatic oil contains at least about 60 weight percent of naphthalenic aromatics and not more than about 20 weight percent of non-aromatics and the catalyst is chromia-alumina.

3. The method of producing naphthalene and high octane aromatic gasoline which comprises contacting in a first stage in the presence of hydrogen a petroleum light cycle oil consisting essentially of aromatics and about 40 to 65 percent of non-aromatic components with a hydroforming catalyst having a non-combustible base at a temperature of about 900 to 1200 F., a space velocity of about 0.1 to 20 LHSV and a pressure of at least about atmospheric while efiecting conversion to high octane aromatic gasoline and an aromatic oil boiling above 400 F. and containing at least about 40 weight percent of naphthalenic aromatics and not more than about 30 weight percent of non-aromatics, said aromatic oil containing not more than about 10 weight percent of naphthalene, separating the high octane aromatic gasoline from the aromatic oil, contacting in a second stage under more severe reaction conditions the aromatic oil With a hydroforming catalyst having a non-combustible base, in the presence of hydrogen and at a temperature of about 900 to 1200 F., a space velocity of about 0.1 to 3 LHSV and a pressure of at least about atmospheric to form naphthalene and an additional quantity of aro matic gasoline, and separating the gasoline and naphthalene.

4. The method of claim 3 in which the aromatic oil contains at least about 60 weight percent of naphthalenic aromatics and not more than about 20 weight percent of non-aromatics and the catalyst is chromia-alumina.

5. The method of claim 3 in which the catalyst employed in both reaction stages is chromia-alumina in which the severity in operating conditions for the first stage is defined by 1000 to 1100 F., 2 LHSV and 200 to 500 p.s.i.g. hydrogen partial pressure and in which the severity in operating conditions in the second stage is defined by 1050 to 1150 F., 0.5 LHSV and 200 to 500 p.s.i.g. hydrogen partial pressure.

6. The method of claim 3 in which the catalyst employed in both reaction stages is molybdena-alumina, in which the severity in operating conditions for the first stage is defined by 900 to 1000 F., 2 LHSV, and to 400 p.s.i.g. hydrogen partial pressure and in which the severity in operating conditions in the second stage is defined by 900 to 1000 F., 1.0 LHSV and 200 to 500 p.s.i.g. hydrogen partial pressure.

7. The method of claim 3 in which beta-methylnaphthalene is also separated from the product of the second stage and the remaining oil boiling above about 400 F. is recycled to the second stage reaction.

8. The method of claim 3 in which 2,6-dimethyl-naphthalene is also separated from the product of the second stage, and the remaining oil boiling above about 400 F. is recycled to the second stage reaction.

9. The method of claim 3 in which the product of the second stage boiling above about 400 F. from which the naphthalene has been recovered, is recycled to the second stage reaction.

References Cited in the file of this patent UNITED STATES PATENTS 2,328,828 Marschner Sept. 7, 1943 2,431,515 Shepardson Nov. 25, 1947 2,653,176 Beckberger Sept. 22, 1953 2,729,688 Anderson et al. Jan. 3, 1956 2,758,062 Arundale et a1 Aug. 7, 1956 2,769,769 Tyson Nov. 6, 1956 2,780,661 Hemminger et a1. Feb. 5, 1957

Claims

1. THE METHOD OF PRODUCING NAPTHALENE AND HIGH OCTANE AROMATIC GSOLINE WHICH COMPRISES CONTACTING IN A FIRST STAGE IN THE PRESENCE OF HYDROGEN A PETROLEUM LIGHT CYCLE OIL CONSISTING ESSENTIALLY OF AROMATICS AND ABOUT 40 TO 65 PERCENT OF NON-AROMATICS COMPONENTS WITH A HYDROFORMING CATALYST HAVING A NON-COMBUSTIBLE BASE AT A TEMPERATURE OF ABOUT 900 TO 1200*F., A SPACE WHEREBY OF ABOUT 0.1 TO 20 LHSV AND A PRESSURE OF AT LEAST ABOUT ATMOSPHERIC WHILE EFFECTING CONVERSION TO AN AROMATIC OIL BOILING ABOVE 400*F. AND CONTAINING AT LEAST ABOUT 40 WEIGHT PERCENT OF NAPTHALENIC AROMATICS AND NOT MORE THAN ABOUT 30 WEIGHT PERCENT OF NON-AROMATICS, SAID AROMATIC OIL CONTAINING NOT MORE THAN ABOUT 10 WEIGHT PERCENT OF NAPHTHALENE, CONTACTING IN A SECOND STAGE UNDER MORE SEVERE REACTION CONDITIONS THE AROMATIC OIL WITH A HYDROFORMING CATALYST HAVING A NON-COMBUSTIBLE BASE, IN THE PRESENCE OF HYDROGEN AND AT A TEMPERATURE OF ABOUT 900 TO 1200*F., A SPACE VELOCITY OF ABOUT 0.1 TO 3 LHSV AND A PRESSURE OF AT LEAST ABOUT ATMOS-