|Publication number||US3159565 A|
|Publication date||Dec 1, 1964|
|Filing date||Sep 26, 1961|
|Priority date||Sep 26, 1961|
|Publication number||US 3159565 A, US 3159565A, US-A-3159565, US3159565 A, US3159565A|
|Inventors||Arey Jr William Floyd, Kimberlin Jr Charles Newton|
|Original Assignee||Exxon Research Engineering Co|
|Export Citation||BiBTeX, EndNote, RefMan|
|Patent Citations (6), Referenced by (4), Classifications (6)|
|External Links: USPTO, USPTO Assignment, Espacenet|
Dec. l, 1964 K|MBER}| |N, JR., ETAL 3,159,565
N. HYDRocAREoN CONVERSION PEocEss To OBTAIN GAsoLINE WITH THE usE oF A SINGLE DIsTILLATIoN ZONE Filed sept. 26. 1961 iChurlesfNewfon Kimberlin, JI. WilfliumfFloyd Arey; Jr. Inventors Paref Attorney bed hydrocracking step. The hydroforming United States Patent O M' 3,159,565 HYDRCARBN QQNVERSIN PRCESS T@ OBTAIN GASQLINE WETH THE USE F A lNGLE DHSTILLATEGN ZGNE Charles Newton Kimhei'lin, lr., and William Fioyd Arcy, Jr., Baton Rouge, La., assignors to Esso Research and Engineering Company, a corporation of Delaware Filed Sept. 26, 196i, Ser. No. ldtl l0 Claims. (Cl. 298-79) This invention relates to a process for converting hydrocarbons such as crude hydrocarbon oil to gasoline. or
The naphtha from the distillation step is hydroforrned and in this step hydrogen is produced and is used in the other steps of the process such as hydrocracking and hydrogenation so that no extraneous hydrogen is needed for the process. With some crudes no extraneous hydrogen is required.
The liquid products from the hydrogenation step or coking step and liquid hydrocracked products are returned to the distillation step so that only one distillation tower is required. The liquid products from the hydroforming step are rerun in a stabilizer tower and bottoms from this stabilizer tower are returned to the single distillation tower. Y
The coking may be a liuid coking step or delayed coking. The hydrocracking step may bea iiuid or fixed or moving may be a iluid step, fixed bed or moving bed operation.
The liquid products from hydrocracking and hydrogenation or coking are returnedto the single distillation tower where the converted hydrocarbons and virgin hydrocarbons are fractionated and the selected fractions then sent to hydroforming, hydrocraclring and hydrogenation or coking to produce primarily a gasoline product to which is added the light naphtha fraction from the single distillation tower. The one distillation tower instead of a number of small fractionating towers serves to fractionate many feeds and provides lower apparatus investment and a saving in operating expense.
Where platinum-on-alumina having combined halogen is the catalyst used in hydroforming, it is preferred to hydroiine the hydrocarbon fraction being fed to the hydroforrner to reduce the sulfur in the feed and minimize contamination and deactivation of the platinum catalyst.
To maintain the hydrogen concentration relatively high in the recycle gas stream it is preferred to pass the hydrogen-containing gas through a conventional absorption step to remove gaseous hydrocarbons and also to desulfurize the recycle gas, if necessary.
According to the present invention an improved combination process is provided in which whole crude petroleumvoil is distilled in a combination tower into a napntha fraction, a middle distillate fraction and a bottoms fraction. The naphtha fraction is catalytically hydroformed and the hydrogen produced is recycled to hydroforming and also is used to supply hydrogen for catalytically hydrocracking the middle distillate fraction. In one form of the invention, the bottoms fraction is coked and the liquid products of coking along with liquid hydrocraclcer products and bottoms from the rerun tower for hydroforrned products are returned to the combination tower.
3,l5,5t5 ralenti-s nee. i, ia64 With this process the crude oil is converted primarily to gasoline. Solid coke is recovered from the coking step. ln another form of the invention the bottoms are hydrogenated under pressure.
The hydrocarbon feed to the hydroformer is a mixture of virgin, hydrocracker and colcer naphtha. The feed to the hydrocracker is a mixture of hydrocracker cycle oil, virgin and coker middle distillates. The feed to the Coker or high pressure hydrogenation step is the liquid hydrocarbon material boiling above that sent to the hydrocracker.
Gasoline is the major product of the process but some heating oil may be withdrawn from the process, if desired. Coke and gases such as C1-C3 hydrocarbon gases are also produced. With certain crude oils, excess butanes will be produced. Hydrogen for the hydrocracking step is obtained from the hydroforming step. There is a transfer of hydrogen between hydrocarbonswithin the process. Where a hydrogenation step is selected for converting the bottoms from the distillation step, the hydrogen is also obtained from the hydroiorming step. Where insuiiicient hydrogen is produced in the hydroforrning step for use in hydrocracking and hydrogenation, extraneous hydrogen is used.
In the drawing the figure represents a iiow diagram including the various steps in the process of the present invention.
Referring now to the drawing, the reference character l@ designates a line for feeding a whole petroleum crude oil into the lower portion of a single distillation or combination tower l2 which may be an atmospheric pressureI distillation zone with provision for suitable fractionation. if desired, a combination atmospheric-vacuum distillation tower or vessel may be used. The whole crude oils which may be used include South Louisiana, West Texas, Mid-Continent, domestic crudes or Middle East crudes, such as Arabian or Kuwait.
A gaseous fraction including CIL-C3 gaseous hydrocarbons is taken off overhead from the tower 12 through line M. This can be used as fuel or may be employed as a source of supplemental hydrogen, if desired.
A bottoms or residual fraction boiling above about G-1000" F. is withdrawn from the bottom of tower l2 through line 16 and in one form of the invention is l passed to coking zone 13 which is preferably a iluid colf.- ing unit but which may be a delayed coking zone. Coke is withdrawn as a solid by-product (not shown). Instead of a coking zone, a high pressure hydrogenation step may be used at 18 for hydrogenating and converting the residual fraction from line 16 to lower boiling and more saturated hydrocarbons.
The converted hydrocarbons from zone 18 are passed through line 2li to a gas-liquid separator 22. for separating gases from liquid hydrocarbons. The gas is taken overhead through line 24 and discarded from the system through line 26 or recovered as fuel gas. Where Zone 18 is a hydrogenation zone, the gas in overhead line 24 is recycled at least in part through line 27 to line 16 for passage and recirculation through the hydrogenation zone i3. Additional hydrogen from line 23 from the hydroforming step later to be described is introduced into line 27 for passage through the hydrogenation Zone i8.'
The liquid hydrocarbons are withdrawn fromthe bottom of the gas-liquid separator 22 and passed through lines Sil and 32 for return to the distillation vessel 12. ln some cases a portion of the liquid hydrocarbons from line 3l) may be passed through line 34 for introduction in the hydrocracking Zone 36 presently to be dcribed.
Where the zone 18 -is a colcing zone, and preferably a uid coking Zone, the coiring step is carried out at a temperature between about 850 F. and 1200 F. at a pres- Vsure between about atmospheric and 15 p.s.i.g. The coke particles are of an average size of between about 100 and 1000 microns. The superficial velocity of the upiiowing gas and vapor in the coking zone is between about 0.2 and 5.0 ft./sec. to maintain a uidized bed of coke particles. The coke circulation rate of solids to oil feed is between about and 10. The burner vessel of the coking unit is maintined at a temperature between about 1050 F. and 1600 F.
Where the zone 18 is a hydrogenation zone, the conversion of the residual oil or bottoms fraction to lower boiling hydrocarbons in zone 18 may be in the range of between about 20 and 80%. The hydrogenation zone contains a fixed bed of catalyst which may be of the nonregenerable type such as M082, W52, or NiSWSz and the pressure in the zone 18 is maintained in the range between about 1500 and 5000 p.s.i., preferably between about 2,000 and 3,000 p.s.i. Instead of using a nonregenerable type of catalyst, a regenerable catalyst may be used such as cobalt molybdate =on silica stabilized alumina in which case the pressure is maintained at about 500 to 1500 psi., preferably between about 800 and 1000 psi. and the catalyst is regenerated periodically at intervals of 1 to 6 months by burning ofr' the deposited carbon. Hydrogencontaining gas (50 to 85% H2) is circulated through the zone 18 from line 27 at a rate between about 1000 and 5000 cubic feet per barrel, preferably about 2000 to 3000 cubic feet per barrel of feed to zone 18. Makeup hydrogen from the hydroforming step presently to be described is introduced into zone 18 from line 23 and some gaseous products are purged from the process through line 26. The oil feed rate to hydrogenation zone 18 is in the range between about 0.5 to 2 w./hr./w. and the temperature is maintained in the range between about 700 F. and 850 F.
Returning to the distillation vessel 12, a light naphtha fraction contining C4 to 160 F.-220 F. virgin and cracked hydrocarbons is taken olf from the upper portion of the vessel 12 through line 37 and used for blending with the gasoline product withdrawn through line 38 from stabilizer tower 40 to be described hereinafter in greater detail.
Further down in the distillation vessel a higher boiling fraction comprising a heavy naphtha fraction is collected and withdrawn through line 42 and passed into the hydroforming zone 44. The catalytic hydroforming zone 44 preferably contains a lixed bed of catalyst employing a platinum on alumina catalyst. The catalyst may contain between about 0.01 and 5 wt. percent platinum on alumina containing between about 0.0 and 5.0 wt. percent of silica. Perferably, the catalyst contains combined halogen in an amount between about 0.3 and 2 wt. percent of fiuorine or chlorine. The temperature of the hydroforming zone is between about 850 and 1000 F., the pressure is between about 50 and 1000 p.s.i.g. and space velocity or v./hr./v. -is between about 0.1 and 10. Hydrogen is produced during hydroforming and some of the hydrogen is passed to the hydrocracking zone to be described hereinafter. Hydrogen-containing gas is recycled to zone 44 through line 46 and contains between about 70 and 90% of hydrogen, and the amount of hydrogen introduced into zone 44 is between about 1000 and 10,000 cubic feet per barrel of naphtha feed. When using a platinum `on aluminum catalyst, it is usually necessary to treat the heavy naphtha feed going to the reforming zone 44 and here a conventional hydrofining process operating at about the same pressure as the reformer but using a cobalt-molybdate-alumina catalyst and a temperature of about 600-750 F. may be used for the removal of sulfur from the heavy naphtha feed. The heavy naphtha fraction boils in the range of about 160- 220 initial and a 300 to 450 F. end point.
The hydroformed products are passed through line 4S to a liquid-gas separator S0 for separating hydrogencontaining gas from liquid hydrocarbons. The gas is passed overhead through line 46 and recycled to the hydroforrning zone 44. The gas in line 46 is preferably passed through a conventional absorption step or the like to remove impurities such as sulfur and nitrogen. The liquid hydrocarbons are withdrawn from the bottom of the separator 50 and passed through line 52 to the stabilizer or rerun tower 40 from which gasoline product is withdrawn through line 30. A gaseous fraction comprising C1-C3 hydrocarbons is withdrawn overhead through line 54 from the tower 40. Higher boiling hydroformed products boiling above the gasoline boiling range are withdrawn from the bottom of rerun tower 40 through line 56 and returned to the single distillation vessel 12.
Instead of having a xed catalyst bed, the hydroforming unit of zone 44 may comprise a fluid reforming zone using a catalyst comprising molybdenum trioXide-onalumina. The conditions employed are 50 to 300 p.s.i.g., temperature of 850-1000 F. and a feed rate of 0.5 to 2 w./hr./w. with a hydrogen recirculation rate of 1000- 5000 s.c.f./bb1.
During the hydroforming, there is a net production of hydrogen and this hydrogen is used to provide for the hydrogen consumption that occurs in the hydrocracking zone 36 and the hydrogenation zone 18.
Returning to the distillation vessel 12, another fraction higher boiling than the heavy naphtha fraction is withdrawn from an intermediate portion of the vessel 12 through line 60 and this fraction has a boiling range between about 300" and 450 F. initial to about 650 F. and `comprises a heating oil or middle distillate fuel product which may be withdrawn in part through line 62 from the process if it is desired to withdraw a heating oil fraction as a product. Instead of withdrawing this fraction as a heating oil fraction, it may be passed through line 64 and line 66 for passage through the hydrocracking zone 36. Further down in the distillation vessel 12 a heavy gas oil fraction, higher boiling than the middle distillate fraction, is withdrawn through line 63 and passed through line 66 to the hydrocraclcing zone 36. With most crude petroleum oils used in this invention, hydrotining of the feed going to the hydrocracker 36 is not necessary. Some of the hydrogen-containing gas from the hydroforming step and from line 46 is passed through line 70 for supplying hydrogen to the hydrocracking step in zone 36.
The hydrocracked products are passed through line '72 into a gas-liquid separator 74 for separating gas from liquid hydrocarbons. The gas being recycled through overhead line '76 to the hydrocracking zone 36. The liquid hydrocarbons withdrawn from the bottom of separator 74 are passed through line 78 for passage through lines 56 and 32 and for return to the single distillation vessel 12. If necessary, extraneous hydrogen can be introduced into line '76 through line S0.
The hydrocracking zone 36 may be a fluid catalyst hydrocracking zone or a fixed bed catalyst hydrocracking zone. In either case with the present invention it is not necessary to hydrone the feed going to the hydrocracking zone 36 and hydrocarbon oil feeds containing more than 50 parts per million of nitrogen can be tolerated. Where a fluid hydrocracking zone is used, the catalyst comprises between 1 and 15% by weight of nickel on silica-alumina cracking catalyst preferably about 6 wt. percent nickel on silica alumina cracking catalyst containing -90% silica and 10-30% alumina. However, higher alumina cracking catalysts may be used if desired. Provision is made for continuous catalyst regeneration by burning with air.
Gther catalysts, such as cobalt on silica-alumina or 0.1-2 wt. percent of platinum or palladium on silicaalumina may be used. The pressure in the hydrocracking zone 36 is between about 200 and 1000 p.s.i.g., preferably between about 400 and 600 p.s.i.g. The temperature in the hydrocracking zone 36 is between about 580 and 700 F., preferably between about 600 and 650 F. Hydrogen or hydrogen-containing gas between about 3,000 and 15,000 cubic feet per barrel, preferably between about and hours, preferably between 1 and 3 hours.
4000 and 6000 cubic feet per barrel of hydrocracking feed is circulated through the hydrocracldng zone 36 through lines 76 and 66. The hydrocarbon feed rate to the hydrocracking zone 36 is between about 0.5 and 5 w./hr./W., preferably 1-2 W./hr./w. and the catalyst holding time in the zone 36 is in the range between about 30 minutes The conversion of hydrocarbons in hydrocracldng zone 36 to products boiling below about 400 F. is rin the range of 30 to about 90%.
Where a fixed catalyst `bed is used in the hydrocracking zone 36 the catalyst contains a sieve-based catalyst, such as palladium on a crystalline zeolitic metal aluminosilicate molecular sieve having uniform pore openings between about 6 and 15 A. units or palladium on a decationized crystalline Vzeolitic metm alumino-silicate molecular sieve having uniform pore openings between about 6 and 15 Angstrorn units. The molecular sieve is furtherA characterized in that it contains no more than about 10% sodium calculated as NagO, preferably between about 0.5 and 8.5%.
The molecular sieves are crystalline zeolitic aluminosilicates that are characterized by having pores of a uniform diarneter. They may be either natural zeo-lites such as Faujasiteor they may be synthetic zeolites such as the 13X or 13Y molecular sieves. The synthetic zeolites are prepared by heat soaking at a temperature of 175 F. to 250 F. for a' period of thirty minutes to several days, a mixture obtainedby mixing a solution of sodium aluminate and a solution of sodium silicate or a silica hydrosol. For example, lSY sieve may be prepared by heat soaking at 210 F. for six days, a mixture obtained by mixing 40.6 pounds of water, 1440 grams of sodium aluminate (46.6% A1203), 5148 grams of sodium hydroxide, and 76.9 pounds of silica hydrosol.v The crystalline product is separated from the mother liquor by filtering and washing with water. The product comprises the sodium form of lSY sieve. The sodium iSY sieve may be converted to the ammonium 13Y sieve by ion exchange with a solution of ammonium chloride. The ammonium sieve may be impregnated with palladium by treatment with a solution of palladium chloride. The palladium impregnated ammonium `sieve is converted to the active catalyst by heating to a temperature in the range of 600 F. to 1000 F. to drive off the ammonia and reduce the palladium to a zero valence state. The amount of the palladium in the catalyst is between 0.01 and 5.0 wt. percent.
The hydrocracking is carried out at a vtemperature between about 300 and 900 F., at a pressure between about 200 and 2000 p.s.i.g. and while recycling between about 3000 and 30,000 standard cubic feet of hydrogen per barrel of hydrocracking feed.
instead of passing all of the hydrogenated or coker products from the bottom of the separator 22 tothe single distillation vessel 12, at least a portion thereof may be recycled through line 34. t-o the hydrocracking zone 36, to increase the saturation of the heating oil fraction drawn off by line 62.
From the vabove description it will be apparent that the hydrofonned products are passed to the stabilizer or rerun tower d0 from which high octane gasoline is withdrawn The rest of the converted v through line 3S as product. liquid products from the hydrocrack-ing step and the coking or hydrogenation step are passed together with high boiling hydroformed products to the single distillation vessel 12 wherein the products are fractionated and the separate fractions are treated as` above described. That is, the fraction withdrawn from the distillation Vessel 12 through line 42 isl a mixture and contains virgin, hydrocracker and Coker naphtha. These are excellent feeds for the hydroforming step. The feed to the hydrocracking zone 36 which is withdrawn from the single distillation Vessel l2 through lines 60, 64 and 68 is a mixture and contains hydrocracker cycle oil, virgin, and coker middle distillates. The residual fraction withdrawn from the bottom of the distillation vessel 212 contains heavy cracked materials and also high boiling residual virgin oil.
By converting all of the fuel oil or middle oil distillate and utilizing the 4hydrogenation zone 1S, it is possible to Vcarry out the process to obtain only high octane gasoline having a research octane number between about and as a liquid product in an amount between about 90 and on the crude oil feed in line l0.
in a specic example, about 10,000 barrels per calendar day of whole crude petroleum oil are introduced or passed through line 10 into the distillation vessel i2. The crude oil of West Texas origin has an API gravity of 31.8. About 9,660 barrels per calendar day of gasoline having a research octane number of 95 unleaded are withdrawn through line 30. About 2,000 barrels per day of a bottoms fraction having a boiling point above about 1000u F. are withdrawn from the vessel l?. through line 16 and fand thisfraction contains virgin residual oil and converted heavy hydrocarbons. 'This bottoms fraction is passed vthrough iluid coliing zone 18 .where the temperature in the coking zone is about 1000 F., the pressure is about 15 p.s.i.g. Ito produce about 1,700 barrels per day of ooker distillate which is passed through lines 30 and 32 back to the distillation vessel l2. The Coker distillate has an initial boiling point of about 32 F. and a final boiling point of about 1000 F.
When cracking all the middle distillate fraction, about 9,000 barrels per day of middle oil distillate having a boiling range between about 430 F. and 1000 F. are passed through lines 60, 64, 68 and 66 to the fixed bed hydro- `cracking zone drnaintained at ya temperature of yabout 650 F. and a pressure of 1500 p.s.i.g. together with recycle gas containing 85% of hydrogen so that about 10,000
cubic feet of hydrogen are supplied per barrel of feed in ing a boiling range between about 180 F. and 430 F.
are Withdrawn from the upper portion of the distillation vessel 12 through line 4t2 and passed through the hydroforming zone 44 which contains a iixed bed of catalyst containing about 0.76% of platinum and 0.6% chlorine on alumina. The hydroforming zone 44 is maintained at a pressure of about 200 p.s.i.g. and at a temperature of about 930 F. The liquid space velocity is about 2 to 3 v./hr./v. About 6,715 barrels per day of hydroformed liquid products are passed through line 52 to the rerun tower 40 for separating 6,580 barrels per day of high octane gasoline from barrels per day of high boiling constituents which are returned to the single distillation vessel 12 through line S6. About 2,000 cubic feet of hydrogencontaining gas 85% H2 per barrel of heavy naphtha feed are introduced into the hydroforming zone 44 through line 46.
The gaseous recycle stream passing through line 46 in the hydroforming step contains impurities such as sulfur and preferably the gaseous stream is passed through a conventional sulfur removal step such as an absorption step n before being recycled through the hydroforming zone Li4 7 rels per day and is blended with the gasoline in line 38 to make the 9,600 barrels total.
Fuel gas in the amount of about million cubic feet per day is also collected as a product from overhead line 14 from distillation tower 12 and line 54 from stabilizer 40. In addition about 53 tons of coke a day are obtained from the fluid coking step.
In the specific example above given the hydroforming step produced enough hydrogen for the hydrocracking and hydroforming step without having to use extraneous hydrogen.
Instead of using a coking step at 18 a hydrogenation step may be used where the catalyst is cobalt molybdatealumina and the pressure in zone 18 is maintained at 1,000 p.s.i.g., the temperature is about 780 F. and hydrogen is supplied at a rate of about 3,500 c.f./b. The oil feed rate is about 0.5 w./hr./w. In this case the amount of hydrogenated hydrocarbon liquid passing through line to tower 12 is about 3,000 barrels per day and the liquid being returned will have about 50% boiling below about 1,000 F. In this case, since very little of the 1,000 F.-} material is converted to coke, there will be a corresponding increase of about 5% in the amount of oil being fed to the hydrocracker and hydroformer as compared to the above case in which the 1,000" F. material was fed to a coker.
In the above examples where the zone 18 is used as a coker or as a hydrogenation zone, no extraneous hydrogen was needed for the process. VThe choice between the coker step and hydrogenation step will depend upon the economic situation at the time, the crude petroleum oil being processed, the demand for coke, etc.
What is claimed is:
1. A process for converting Whole petroleum crude oil mostly to gasoline which comprises distilling whole petroleum crude oil in a single distillation Zone into a bottoms fraction boiling above about 1000 F., a heavy naphtha fraction and a middle distillate fraction higher boiling than said heavy naphtha fraction, hydroforming said heavy naphtha fraction and separating a gasoline fraction from higher .boiling hydroformed liquid hydrocarbons, catalytically hydrocracking said middle distillate fraction, thermally converting said bottoms fraction to lower boiling hydrocarbons, returning all of said separated higher boiling hydroformed liquid hydrocarbons and all of said liquid hydrocracked products and all of said converted liquid bottoms fraction to said single distillation Zone for fractionation so that converted products are fractionated along with virgin hydrocarbons to provide an improved feed for said hydroforming step for producing additional hydroformed gasoline.
2. A process according to claim 1 wherein said bottoms conversion step comprises a coking step.
3. A process according to claim l wherein said bottoms conversion step comprises a high pressure hydrogenation step.
4. A process according to claim 2 wherein said feed to said hydroforming step comprises a mixture of virgin naphtha, coker naphtha and hydrocracker naphtha and said feed to said hydrocracking step comprises a mixture of hydrocracker cycle oil, virgin middle distillate fraction and coker middle distillate fraction.
5. A process for converting Whole petroleum crude oil into lower boiling hydrocarbons in a combination process by hydrogen transferwithin said process to produce lower boiling hydrocarbons including about the same volume of gasoline as the crude oil feed without the need of supplying extraneous hydrogen which comprises distilling whole petroleum crude oil in a single distillation zone in the presence of cracked or converted hydrocarbons produced as hereinafter set forth to separate a bottoms residual fraction containing virgin and cracked hydrocarbons and boiling above about 1000 F., a heavy naphtha fraction containing virgin and cracked hydrocarbons and a middle distillate fraction higher boiling than said heavy naphtha and containing virgin and cracked hydrocarbons, hydroforming said heavy naphtha fraction and separating a gasoline fraction from higher boiling hydroforrned liquid hydrocarbons and from hydrogen produced during hydroforming, recycling a portion of said separated hydrogen to said hydroforming step, said hydroi'ormed gasoline forming the major product of the process, catalytically hydrocracking said middle distillate fraction in the presence of recycled hydrogen produced in said hydroforming step, separating hydrogen-containing gas from the hydrocracked liquid products, recycling said hydrogen-containing gas to said hydrocracking step, converting said bottoms residual fraction to lower boiiing hydrocarbons and returning all of said higher boiling hydroformed liquid hydrocarbons and all of saidy liquid hydrocracked products and all of said liquid converted bottoms residual fraction to said single distillation zone for fractionation into fractions as above set forth to provide an improved hydrocarbon feed for said hydroforming step.
6. A process according to claim 5 wherein a light naphtha fraction containing cracked and virgin constituents is withdrawn from the upper portion of said distillation zone and is blended with the gasoline recovered from said hydroforming step.
7. A process for converting whole petroleum crude oil mostly to gasoline which comprises distilling whole petroleum crude oil in a single distillation zone in the presence of cracked or converted hydrocarbons produced as hereinafter set forth to separate a bottoms residual fraction containing virgin and converted hydrocarbons, a heavy naphtha fraction containing virgin and converted hydrocarbons and a middle distillate fraction higher boiling than said heavy naphtha fraction and containing virgin and converted hydrocarbons, hydroforming said heavy naphtha fraction and separating from higher boiling hydroformed liquid hydrocarbons and from hydrogen produced during hydroforming a gasoline fraction which forms the major product of the process, catalytically hydrocracking said middle distillate fraction in the presence of hydrogen produced in said hydroforming step, separating hydrogencontaining gas from the liquid hydrocracked products, recycling said hydrogen-containing gas to said hydrocracking step, converting said bottoms residual fraction to lower boiling hydrocarbons and returning all of said higher boiling hydroformed liquid hydrocarbons and all of said liquid hydrocracked products and all of said liquid hydrocarbons from said converted bottoms residual fraction to said single distillation zone for fractionation into said fractions above set forth to provide an improved feed for the hydroforrning step.
8. A process according to claim 7 wherein said bottoms fraction from said single distillation zone is coked and the liquid Coker products are passed to said single distillation Zone to be fractionated.
9. A process according to claim 7 wherein said bottoms fraction from said single distillation zone is coked and the liquid coker products are at least in part recycled to said hydrocracking step.
10. A process for converting whole petroleum crude oil into lower boiling hydrocarbons in a combination process by hydrogen transfer within said process to produce lower boiling hydrocarbons including gasoline produced by hydroforming in an amount between about and 110% on the whole crude oil feed which comprises distilling whole petroleum crude oil in a single distillation zone in the presence of cracked or converted hydrocarbons produced as hereinafter set forth to separate a bottoms residual fraction containing virgin and cracked hydrocarbons and boiling above about 1000 F., a heavy naphtha fraction containing virgin and cracked hydrocarbons and a middle distillate fraction higher boiling than said heavy naphtha and containing virgin and cracked hydrocarbons, hydroforming said heavy naphtha fraction and separating a gasoline fraction from higher boiling hydroformed liquid hydrocarbons and from hydrogen produced during hydroforming, which gasoline forms the major product of the process, catalytically hydrocracking said middle distillate fraction in the presence of hydrogen produced in said hydroforming step, separating hydrogen-containing gas from the hydrocracked liquid products, recycling said hydrogen-containing gas to said hydrocracking step, converting said bottoms residual fraction to lower boiling hydrocarbons and returning all of said higher boiling hydroformed liquid hydroc-arbons and all of said liquid hydrocracked products and all of said converted bottoms residual fraction to said single distillationfzone for fractionation into fractions to be treated as above set forth to provide an improved feed for said hydroforrning step for producing additional hydroformed gasoline.
References Cited in the le of this patent UNITED STATES PATENTS 2,294,126 Ocon Aug. 25, 1942 2,312,445 Ruthruf Mar. 2, 1943 2,731,396 Harding et al Ian. 17, 1956 2,777,801 Bittner et al. Jan. 15, 1957 2,853,439 Ernst Sept. 23, 1958 3,019,180 Schreiner et al lan. 30, 1962
|Cited Patent||Filing date||Publication date||Applicant||Title|
|US2294126 *||Sep 12, 1938||Aug 25, 1942||Ocon Ernest A||Method of treating a plurality of hydrocarbon oils for subsequent cracking|
|US2312445 *||May 13, 1940||Mar 2, 1943||Robert F Ruthruff||Catalytic combination process|
|US2731396 *||Jul 30, 1952||Jan 17, 1956||Exxon Research Engineering Co||Combination crude distillation and cracking process|
|US2777801 *||Dec 3, 1951||Jan 15, 1957||Exxon Research Engineering Co||Combination crude distillation and oil refining process|
|US2853439 *||Aug 1, 1952||Sep 23, 1958||Exxon Research Engineering Co||Combination distillation and hydrocarbon conversion process|
|US3019180 *||Feb 20, 1959||Jan 30, 1962||Socony Mobil Oil Co Inc||Conversion of high boiling hydrocarbons|
|Citing Patent||Filing date||Publication date||Applicant||Title|
|US3787314 *||Nov 21, 1972||Jan 22, 1974||Universal Oil Prod Co||Production of high-octane, unleaded motor fuel|
|US4082647 *||Dec 9, 1976||Apr 4, 1978||Uop Inc.||Simultaneous and continuous hydrocracking production of maximum distillate and optimum lube oil base stock|
|US5958218 *||Jan 22, 1996||Sep 28, 1999||The M. W. Kellogg Company||Two-stage hydroprocessing reaction scheme with series recycle gas flow|
|USRE33323 *||Jun 10, 1988||Sep 4, 1990||Exxon Research & Engineering Company||Reforming process for enhanced benzene yield|
|U.S. Classification||208/79, 208/80, 208/78|