US3172840A - Light ends - Google Patents

Description

N, J. PATERsoN 3,172,840 HYDROCARBON CONVERSION PROCESS 2 Sheets-Sheet 1 March 9, 1965 Filed July 2. 1962 March 9, 1965 N. J. PATERsoN HYDRocARBoN CONVERSION PRocEss 2 Sheets-Sheet 2 Filed July 2, 1962 mXOU jmtjm mmxou uur] EDDnZmmm +.moom`l INVENTOR NORMAN J. PA TERSON ATTORN EYS United States Patent O 3,172,840 HYDROCARBON CONVERSION PROCESS Norman J. Paterson, San Rafael, Calif., assigner to California Research Corporation, San Francisco, Calif., a corporation of Delaware Filed July 2, 1962, Ser. No. 206,762 6 Claims. (Cl. 298-79) INTRODUCTION This invention relates to the conversion of hydrocarbonaceous materials such as petroleum oils, tarsand oils, shale oils and coal-oils, including residual and nonresidual portions thereof, to gasoline and middle distillates. Particularly, the invention relates to an integrated refinery process including units for producing, in high liquid yields, hydrocarbons boiling within the gasoline and middle distillate boiling ranges.

PRIOR ART PROCESSING GENERALLY, AND GASOLINE YIELDS =In conventional refinery operations, the crude oil is passed to a distillation unit normally called a crude column, and the oil is fractionated into various cuts, including light gasoline, heavy gasoline, light and heavy gas oil fractions, and a residual portion boiling so high as t resist vaporization in the crude column. This residual portion is subjected to additional processing, such as vacuurn distillation and coking, in order to produce additional gas oil distillate fractions.

In conventional refinery operations conducted for the purpose of maximizing the production of gasoline, the aforesaid light and heavy gas oils from the crude column, along With the aforesaid heavy gas oil fractions obtained by such further processing of the residual crude oil portion as vacuum distillation and coking, have been used as feed to a conventional catalytic cracker or a thermal cracker,

In these conventional refinery operations, the conversion in the catalytic cracker of the various gas oil fractions to fractions boiling in the gasoline range varies widely. Thus, light straight run gas oil is converted to gasoline at a yield of approximately 40%, and heavy straight run gas oil is converted to gasoline at a yield of approximately 60%, whereas light and heavy Coker gas oils, due to the refractory nature of the same and the presence of catalytic poisons, are converted to gasoline at low yields of about 30% and 25%, respectively.

PROCESSING `PROBLEMS AND FACTORS INFLU- ENCING YIELDS, QUALITY AND STABILITY OF GASOLINE, AND MIDDLE DISTILLATES Although the usual concept is that crude petroleum consists entirely of hydrocarbon components, many petroleums, especially those of `Californie. and Venezuelan origin, have high contents of non-hydrocarbons. For example, Wilmington, California, crude oil contains ten times as much sulfur as Ponca City, Oklahom, crude oil and 65 times the amount of nitrogen. Crude oil from Boscan, Venezuela, compared to Baxterville, Mississippi, crude of the same gravity and Conradson carbon content contains about 3() times the amount of vanadium. Many of the difiiculties of processing and of preparing satisfactory products from petroleum have been attributed to the non-hydrocarbon components. Among the problems of processing, those of catalyst poisoning in catalytic cracking by nitrogen and metallic constituents, are well known. In subsequent processing, sulfur, nitrogen and metallic constituents affect product yields, quality and product stability. Generally, the non-hydrocarbon portions of petroleum tend to concentrate in the heaviest fractions of the crude oil. Crude oils vary in their content of asphaltic residue r'ce dal

which, in turn, is generally a function of the Conradson carbon content of the topped crude oil after removing the 400 F. end point gasoline. Thus, Arabian crude, 34 API gravity, has a Conradson carbon content on residue after gasoline of 4.4 Weight percent. This compares to Boscan, Venezuela, crude oil of 11 API gravity, which has a Conradson carbon of 15 Weight percent on the gasoline-free residue. Thus, the Conradson carbon residue test is a function of the asphaltene content of the crude Y PRIOR ART METHODS OF OBTAINING ADDI- TIONAL CATALYTIC CRACKING FEED STOCKS lFROM STRAIGHT RUN RESIDUAL FRACTIONS Coking, solvent deasphalting and high pressure hydrogenation have been suggested as possible means for reducing the yields of straight run residues and converting them in part to additional gas oil feed stocks for catalytic cracklng.

Hydrocracking of residual stocks is not practiced at present, due .to the high pressures required and high hydrogen consumptions involved. Solvent deasphalting is an effective feed preparation process for producing additional feed stocks for catalytic cracking. With low asphalt or low Conradson carbon content crudes, there is considerable justification for utilizing this process since the ratio of additional cracker feed to the amount of asphalt produced is high. F or high asphaltic content crude oils, however, coking has certain advantages that make it attractive for reducing straight run residues. For example, it is not critical with respect to feed stocks since there is no practical limitation on ash, sulfur, nitrogen, metals content, or Conradson carbon residue. Since there are no limitations on crude source, the application of coking involves the minimum of feed preparation which tends to reduce operating costs. Variations in product distribution can be obtained by varying recycle cut points and rates and combining coking with catalytic or thermal cracking, From;

the standpoint of straight run residues, coking competes with solvent deasphalting except Where complete eliminar. tion of asphalt is desired. For catalytic cracking feed preparation, high asphaltic residues are best handled in a coking unit because higher yields of cracker feed are produced compared to solvent deasphalting. In addition, Where the heavy asphaltic residue constitutes a high 'per. centage of the crude oil, and generally it follows that the metallic constituents are also high, it is preferable to handle this type of feed stock in a coker to prepare low metal content gas oils for subsequent cracking. However, the selectivity of the coking operation is decidedly inferior to catalytic cracking; and, in order to obtain maximum yields of finished products, it is desirable to maximize production of gas oils in the coking unit as feed to the catalytic cracker and obtain the higher product selectivity in the catalytic cracking process.

PRIOR ART USE OF HYDROCRACKING INSTEAD OF CATALYTIC CRACKING TO CO'NVER STRAIGHT RUN GAS OILS Recently, hydrocracking has been practiced to convert gas oils to high yields of gasoline and middle distillates. Since this process exhibits higher product selectivity than catalytic cracking, the selectivity of the hydrocracking Patented Mar. 9, 1965 process to produce high yields of liquid products, such as gasoline and middle distillates, is affected by the presence of nitrogen-containing compounds within the hydrocracking charge stock boiling range. These nitrogenous compounds, examples of which include pyrroles, amines, indoles and other classifications of organic compounds, result in the deactivation of the catalytically active metallic components, as well as the refractory inorganic oxide carrier material which acts as the acidic component of the hydrocracking catalyst. Such deactivation appears to result lthrough the reaction of a nitrogenous compound with the various catalytic components, the extent of such deactivation steadily increasing as the process continues and as the nitrogen-containing feed stocks continue to conltaminate the catalyst through contact therewith. The more predominating eifect of nitrogen deactivation is believed to be the formation of nitrogen-containing complex through interreaction with the catalytically active metallic components, whereby the active centers of the catalyst, normally available to the hydrocarbon charge stock, are effectively shielded therefrom. Deactivation of this nature is not believed to be a simple reversible phenomena which may be easily rectified by merely heating the catalyst in the presence of hydrogen for the purpose of decomposing the nitrogen-containing complexes.

PRIOR ART USE OF HYDROCRACKING INSTEAD OF CATALYTIC CRAOKING TO CONVERT COKER GAS OILS To remove the nitrogen deactivating materials from the feed stock to the hydrocracking process, it is customary to supply the feed stock to an initial hydrotining step. Although the major proportion of the nitrogenous compounds may be removed in this step, it is very dimcult to remove all of the nitrogenous material from the feed stocks boiling in excess of 1000 F. In particular, gas oils produced from high boiling feed stocks such as petroleum residues by thermal means, such as coking, contain nitrogenous materials that are diiiicult to remove if the boiling point of the gas oil is in excess of, say, 750 to 850 F. The causes for increased difficulty of nitrogen removal with increased boiling point are at least two-fold. First, the amount of nitrogen increases markedly with increased boiling range. Secondly, it appears that the increased size of the non-nitrogen portion of the molecule reduces the rate of constant of denitrification, perhaps by` making catalyst contacting more difficult. In addition, the tendency for higher boiling stocks to contain more of the aromatic-type nitrogen compounds results in lower reaction rates. Since hydrogenation of the nitrogen-containing ring and cracking of the ring must precede the final removal of ammonia, the relative inaccessibility of this ring in higher boiling stocks probably contributes to lower reaction rates. Heavy gas oils produced from high nitrogencontaining crude oils, such as Wilmington, California crude, will have total nitrogen contents as high as 6000 p.p.m. for the 800 to l050 F. boiling point gas oil. Coker gas oils having a 750 to 1000 F. boiling range from the same crude will have nitrogen contents as high as 10,000 p.p.m. total nitrogen. To reduce this nitrogen to a level of less than ppm., preferably below 0.5 p.p.m., very severe hydroiining conditions would be required that would not be economic or attainable with present-day hydrolining catalysts. Accordingly, it is not conventional practice to attempt to hydroiine and then hydrocrack'such high boiling vcoker gas oils.

`From the foregoing, it is clear that, to provide feed stock suitable for selective hydrocracking from a high nitrogencontaining crude oil, in order to maximize gasoline from this crude oil, requires the installation in a refinery of a coking process modified to reduce the boiling range of the overall gas oils in order -to reduce the severity of operating conditions in the hydroiining step. It is well known `in the art that two types of Coking processes are available, namely, delayed coking and fluid coking. The type of process selected will depend to a large extent on the characteristics of the feed stock and the type of cracking process available to process the gas oils from the coking step.

DELAYED COKING PROCESSES Delayed coking uses a pipestill heater operating with a maximum of 900 to 950 F. heater outlet temperature in order to avoid coke formation in the heater tubes and transfer lines. There are usually two coke drums connected to a coking heater, one drum being on-stream while the other is being cleaned and prepared for use. The coke drums operate at a lower temperature, usually in the range of 800 to 900 F. because of flashing and endothermic heat of reaction, and the drums provide long residence time favoring the formation of coke and lighter products. The drums usually operate at 40 to 50 p.s.i.g. The superficial vapor velocities in the drum are of the order of 0,2 foot per second and, with a drum height of approximately 40 feet, residence time in the drum is of :the order of 200 seconds, and total residence time, if the heater tubes and transfer lines are included, will be somewhat higher,

FLUlD COKNG PROCESSES Fluid coking differs fundamentally from delayed coking in that contact 'times are much shorter and the reaction .temperatures higher because the necessary feed preheat yand heat of reaction are provided by the circulation of hot Acoke particles from a burner which uses the coke product itself as a fuel. With reactor velocities in the range of 2 feet per second, the residence times are in the order of 15 to 20 secondes at substantially isothermal reaction temp-eratures, normally 950 to 1000 F. Operating pressures in the fluid Coker reactor are usually about l0 p.s.i.g. and theoretically, at least, the lower pressures, compared lto delayed ooking, favor the cracking reaction which is accompanied by a very large increase in the volume of vapor products. As a result of the higher temperatures and short-er residence time in the iluid coker reactor, gas oil fractions (850 to l000 F. boiling range) in 1the feed tend to lliash off rather than crack. The iluid coker recycle stream is generally a l000 F.{- material. Attempts to recycle Coker gas oil, boiling much below a 1000 F. generally do not give' satisfactory resuflts. When it is desired to obtain a lower end point flu-id ook-er ,gas oil, the amount of heavy gas oil recycle will be quite large. With the fluid coker reactor normally 4operating at about 950 F., the heavy gas oil recycle will be converted at this tempenature predominantly in the vapor phase and will be severely degra ed to gas because of the relatively high vapor cracking intensity and because of the high recycle rate. It follows lthen that the liuid coker gas oil usually will have an end point of labout 950 to l000 F. compared to 800 to 950 F. for the delayed coking process.

FLUID COKING VERSUS DELAYED COKING The foregoing considerations indicate that, on a given feed stock to a iuid coker, .the coke production is less and the yield of coker -gas oil is considerably more in a fluid coking process than in a delayed ycoking process, although the gasoline yield is lower and gas production similar. Also, as :a Aresult of the higher reactor temperatures, the products from fluid coking tend to be more oleiinic and the gas oil products more unstable. This follows from the lower coke yield in fluid coking and the lower hydrogen-to-carbon ratios -in the liquid products. Accordingly, it it were desired to lprocess `a low gravity, high nitrogen crude oil for maximum gasoline yield, with processing steps including coking and hydrocracking, it would be preferable to employ the delayed coking process tin order to maximize yields -of -gas oil feed stock of low enough boiling point and nitrogen content as feed for the hydrocracking step. However, it has been found that, even when a conventional single coil delayed Coker is used, the 800 to 950 F. end point of the delayed coker gas oil presents a serious problem. This problem exists because organically-bound nitrogenous compounds in particular are very diicult to remove from the Coker gas oil (thermally cracked) fractions boiling above 750 to 800 F., although virtually complete removal could be effected from the coker gas oil if it had a maximum end point below 750 F. To provide a Conventional single coil delayed coker gas oil having a maximum end point of 75 0 F. from a high nitrogen crude, the 750 E+ long residuum would require extensive recycle of the gas oil fractions through the coking heater and coke drum. Since the delayed coking heater processes a mixture of asphaltic residue along with recycle gas oil and is limited to heater outlet temperatures of about 950 F., a recycle ratio of as hi las 4:1, based on fresh feed, may be encountered. Extensive recycling under these conditions of temperature and pressure would produce large -amounts of gas and degrade the gas oil product due to .the extensive recycling required to produce the desired end point gas oil.

OBJECTS In view of the foregoing, it is an object of the present invention to provide an integrated refinery process wherein residual stocks lare eiciently converted to gas oil feed stock-s for further prccesing.

It is a further object of the present invention .to provide an integrated renery process wherein residual stocks are efficiently converted in part to gas oils which are hydrocnacked to produce good yields of valuable fuel products.

It is a further object of the present Iinvention to p-rovide an improvement in `a process for hydrocracking of Coker gas oils.

DRAWINGS The invention will best be understood, and further objects and advantages thereof will be apparent, from the following detailed description when read in conjunction with the accompanying drawings, in which:

FIG. 1 is a diagrammatic illustration of process units and llow paths of a two-stage delayed coker; and

FIG. 2 is a diagrammatic illustration of process units and ilow paths wherein coking, reforming and `hydro- Cracking operations are combined in an integrated operation.

STATEMENT OF INVENTION In accordance with the embodiment of the present invention, there is provided, in a process for converting a crude hydrocarbon feed having an :end .point above 950 F. to lower boiling valuable liquid products which comprises separating said crude hydrocarbon feed into fractions including lat least one naphtha fraction and .at least one residuum fraction substantially boi-ling above 750 treating ysaid residuum fraction in `a coking zone to produce coke and a liquid Coker distillate, and contacting said liquid Coker distillate in a hydrocracking zone in the presence of hydrogen and a hydrocracking catalyst under hydrocracking conditions, the improvement wherein said coking zone is a two-stage delayed coker comprising a thermal cracking furnace and a Lcoking furnace, and at least one coke drum, and wherein said residuum is treated ytherein by passing said residuum into a coker bubble tower, passing Ia side stream boiling in the range of about 750 to 950 F. from said bubble tower to said thermal cracking furnace, passing the effluent from said thermal cracking furnace to said coke drum, passing la bubble tower residuum stream from said bubble tower to said coking furnace, passing the effluent from said coking furnace to said coke drum, passing the eilluent from said coke drum to said 4bubble tower, and recovering from said bubble tower said liquid coker distillate for further processing in said hydrocracking zone.

e Pnocrss OPERATION Referring now to FIG. 1, there shown is a diagrammatic illustration of an embodiment of process units and flow paths for a twoestage (two-coil) delayed coker suitable for use in practicing the present invention. The twocoil terminology refers to: (l) a 'first coil, or coker furnace, used for heating the feed to the coke drums whereby lthe combination of said Coker coil and coke drums serves to decarbonize said feed; and (2) a second coil, or furnace operating under high pressure, to accomplish boiling point reduction by thermal crack-ing, under selected conditions, of a side stream of particular range from the coker bubble tower.

In the embodiment shown in FIG. l, a crude hydrocarbon feed having an end .point above about 950 F., for example a 20 API Los Angeles Basin crude, is passed through line 1 into distillation column Q. where it is separated into various fractions, including a heavy straight run naphtha boiling from about 200 to 380 F., at least one gas loil boiling between about 380 and 750 F., and a long residuum boiling 'above about 750 F. The various naphtha and gas oil fractions are removed from distillation Column 2, `for example, through lines 3, 4, y5 and 6 as shown. Further processing Iof these fractions will be discussed below in connection with FIG. 2.

lFrom distillation column 2, the long residuum boiling above about 750 F. is passed through line 7 to conventional coker bubble tower 8 as shown. This long residuum picks up heat in bubble tower 8 by direct heat exchange with hot vapors entering bubble tower 8 through line 9. From the bottom of bubble tower 8 la combined feed (residuum plus heavy cycle stock) is passed through -line 10 to coking furnace, or coking coil, 15 where its temperature is raised to about 900 to 950 F. The residuum so heated is passed from cokfing furnace 1S through lines 1.6 and 17 into .the bottom of one of the two coke drums 13A and 18B, each of which is alternately onstreaim while the other is being cleaned and prepared for use.

It is generally the practice to preheat the cleaned empty drum, prior to placing it black on stream to receive the effluent from the coker furnace, by bleeding into it some of the hot overhead vapor from the on-stream drum. However, lthe overhead vapor from the yon-stream drum is only at a temperature of about 800 F. and considerable liquid that condenses in .the drum being preheated must be pumped out and recharged to the coker furnace. Accordingly, considerable time, for example six to eight hours, may be involved in the preheating procedure. In accordance with the present invention, a great deal of time can be saved by preheating the empty cleaned drum with hot effluent from the thermal cracking coil, which provides la higher level of :available heat for preheating. Once Cleaned, .the drum so preheated may be charged with the effluent from the Coker heater together with the thermal coil eflluent.

Operating conditions in coke drums 18A and 18B have been discussed above under delayed coking processes, except that, in order .to maximize the yield of :gas oil, it is preferred to operate the coke drums at a pressure not exceeding 25 p.s.i.\g., and preferably from 5 to 15 p.s.i.g. Overhead vapors from coke drums 18A or 18B are passed through lin-e 9 to bubble tower `8.

A side stream boiling between about 0 and 950 F. is passed from coker bubble tower 3 through line 19 to Aa second coil, or furnace, 20 Where it is thermally cracked at a temperature of between 850 Iand 1000 F. Iand a coil outlet pressure of between 300 and 1000 psig., to effect conversion .to lower boiling products. Conversion in this thermal cracking coil preferably is controlled to maximize yield olf 380 to 750 gas oil and Ito minimize the production of 380 F. end point gasoline.

It is within the scope of the present invention -to introduce extraneous stocks directly into thermal cracleing Coil 20. In such cases thermal cracking coil 20 is an eicient annoio means of reducing the boiling points of such stocks and also reducing the levels of metals, nitrogen and other contaminants thereof. While such extraneous stocks also could be introduced into the feed .to .the ooker furnace, the boiling point reduction thereof and lcontaminant removal therefrom can be accomplished in ya more controlled manner if they are introduced directly into the thermal cracking coil 20.

The effluent from thermal cracking furnace 20, after pressure reduction and quenching (not shown) is passed through lines 25 and 17 to the Aon-stream coke drum 18A or 18B.

From coker bubble tower 8, materials boiling below about 380 F. are passed .through line 2.6 to separator 27. From separator 27, light gases are remo-ved through line 28, and a naphtha fnaction boiling from about C5 to 380 F. is passed through line 29 to a hydrotining zone as shown in FIG. 2. From coker bubble tower 8, coker distillate boiling between labout 380 and 750 is passed .through line 30 toa hydrocracking zone as shown in FIG. 2. In addi-tion to reflux supplied to .the top of coker bubble tower 8, it is sometimes advisable to remove additional heat by circulating an intermediate reflux stream.

From the foregoing description of a two-coil operation, it may be seen that, from a crude oil feed stock having an end point above about 950 F., a C5 to 380 F. naphtha may be produced, as well as a gas oil having an end point of about 750 F., suitable as a feed to a hydrocracking zone, all with an inherent reduction in feed nitrogen content, which is characteristic of the operation of the Coker. It would not be satisfactory to attempt to produce a 380 to 750 F. coker distillate, suitable for hydrocracking, by recycling in a single coil delayed coker; in such a coker, the coil outlet temperatures on the furnace are limited by the coke formation from the heavy asphaltic residue in the tubes; this would necessitate a large furnace and high recycle rates, resulting in an excessive amount of gas production.

Operation of the thermal cracking furnace 20 to give maximum yields of gas oil, as discussed above, may be obtained by controlling conversion in thermal cracking furnace 20 to give 25 to 30 weight percent of 380 F.-conversion products. This conversion may be obtained by controlling crack per pass in furnace 20 by proper selection of top transfer temperature and recycle ratio. As total conversion, measured by weight percent of 380 F.-conversion products increases, the yield of 380 to 750 F. gas oil goes through a maximum. Operation to give 25 to 30 weight percent of 380 F.-conversion products, as discussed above, will insure maximum production of gas oil cut because it is known that when certain combinations of temperature, pressure and reaction time are reached in furnace 20, the gas oil produced will crack faster than it will appear as an end product. Although a yield of 380 to 750 F. gas oil increases with conversion, up to a maximum, gas oil to gasoline ratio always decreases with increasing conversion.

The effluent passing tmough line 25 from thermal cracking furnace 2-0 passes through a pressure reduction valve which reduces the pressure from 300 to 1000 p.s.i.g. to approximately 50 p.s.i.g., prior to the passage of this efiiuent into the bottom of the on-stream coke drum 18A or 18B. The net effect of combining the stream in line 16 from coker furnace 15 with the stream in line 25 from thermal cracking furnace 20 is to supply more heat to coke drum 18A or 18B without the necessity of raising the outlet temperature to coking furnace 15, than if the materials in line 25 were not passed to the coke drums. The additional heat supplied to the coke drums from the heated materials in line 25 permit more gas oil to flash off and also provide a higher level of available heat in the vapors leaving the coke drums through line 9. This in turn improves fractionation in the coker bubble tower and permits better separation of the desired gas oil prod- 8 ucts. If the additional heat for the coke drums that is provided by the heated materials in line 25 were not provided, it would be necessary to raise the temperature of the coker furnace outlet, and coking problems would be encountered in the coking furnace.

Example 1 As an example of the preferred method of operation explained above, a 14 API Los Angeles Basin long residuum having an initial boiling point of approximately 750 F. by ASTM D-ll60 method was processed in a single-coil operation to a 750 F. end point gas oil; and, for comparison, the same operation is shown in the preferred manner of the present invention utilizing a two- From the above it can be noted that there is an increase of 10% in the volume of liquid products with :operation in accordance with preferred operation of the present invention, compared with the single-coil operation. in addition, the amount of feed available for conversion in the subsequent hydrocracking step is increased about 16%. In the thermal cracking coil, conversion of the heavy gas oil boiling above 750 F. takes place predominantly in the liquid phase; and, as a result, yields and qualities of the 380 to 750 F. gas oil are improved. vBased on the total nitrogen content of the 750 F.llong residuum of approximately 9000 ppm., the nitrogen con- .tent of the 380 to 750 F. will be approximately 2000 ppm.

As heretofore mentioned, the contaminants in the gas 'oil feed to the hydrocracking step, particularly organically-bound nitrogenous compounds, are very difficult to remove from thermally cracked fractions boiling above :about 750 to 850 F.; whereas, if the end point is limited to say, 750 F., virtually complete removal of the nitrogen contaminants may be effected. Through utilization of the two-coil delayed coking process followed by a hydrocracking step, in accordance with the present invention, and when processing hydrocarbon feed stocks boiling above about 750 F. to 1000 F. or more, it is possible to produce high volumetric yields of gasoline boiling range hydrocarbons while simultaneously maximizing the yield of middle-distillate hydrocarbons free from nitrogenous compounds. A substantial increase in the overall yield of gasoline boiling range hydrocarbons may then be obtained by further processing nitrogen-free middle-distillate material simultaneously produced by the present process. The fiexibility of the present process permits the withdrawal to storage of the middle-distillate hydrocarbon product for subsequent conversion to gasoline boiling range hydrocarbons when market conditions so dictate.

One of the primary functions to be served by the twostage delayed coking zone is the conversion of those hydrocarbons boiling in excess of a temperature of 750 F. into lower boiling hydrocarbon products which boil below about 750 F. An additional function of the two-stage coking zone is the conversion of materials boiling in excess of a temperature of 750 F. into high yields of gas oil product having a boiling range of over 300 F. initial boiling point and an end point of approximately 750 F. while minimizing the production of materials boiling below 300 F. A further function is the production of maximum yield of gas oil having a nitrogen content in the range of 100 to 3000 p.p.m. from materials boiling in excess of 750 F. having a nitrogen content in excess of 2000 p.p.m. By use of the two-stage delayed coking process, the gas oil produced for subsequent conversion in a hydrocracking zone will not be degraded in quality compared to a similar boiling range gas oil produced in a single-coil coking operation or by other coking processes.

It is understood that the feed to the thermal cracking coil in the present invention will be Coker gas oil boiling in excess of 750 F. produced in the first stage and may be augmented with gas oils, both straight run and thermal and catalytic cycle oils from outside sources.

Referring now to FIG. 2, there shown is a diagrammatic illustration of an embodiment of process units and flow paths, including a two-stage (two-coil) delayed coker, a catalytic reforming zone, a hydrofining zone, and a hydrocracking zone, arranged in an integrated manner in accordance with the process of the present invention.

In the embodiment shown in FIG. 2, a crude hydrocarbon feed having an end point above about 950 F. is passed through line 40, which corresponds to line 1 in FIG. 1, to distillation column 41, which corresponds to distillation column 2 in FIG. l, where it is separated into various fractions as discussed in connection with FIG. l. Light ends and light straight run naphtha are removed from the distillation colunm 4l through line 42, which corresponds to line 3 in FIG. l. A heavy straight run naphtha is passed from distillation column 41 through line 43, which corresponds to line 4 in FIG. 1, to hydrotining zone 44 if nitrogen and sulfur removal is necessary; otherwise, it is passed directly to catalytic reforming zone 45. If nitrogen removal is necessary, the denitried naphtha is passed from hydrofining zone 44 to catalytic reforming zone 45 through line 46.

From distillation column 41, at least one gas oil fraction is passed through line 47, which corresponds to lines 5 and/or 6 in FIG. l, to hydrocracking zone 48. If desired, this fraction first may be denitriiied in the manner discussed below in connection with the coke distillate fraction in line 64.

From distillation column 41 the long residuum boiling above 750 F. is passed through line 49, which corresponds to line '7 in FIG. 1, to delayed Coker 50 which includes coke drums, coking furnace, thermal cracking furnace and coker bubble tower, and which is shown in detail in FIG. 1. Coke that is removed from the coke drums of delayed Coker 50 is indicated as leaving the system by line 60. Light ends are withdrawn from coker 50 through line 61, which corresponds to line 26 in FIG. l.

A C5 to 380 F. naphtha is passed from coker 50 to hydrofining zone 44 through line 62, which corresponds to line 29 in FIG. 1. The oletins in this naphtha are saturated in hydroiining zone 44, after which a portion of the naphtha is passed to catalytic reforming zone 45 through line 46. This olefin saturation is accomplished because the olelins would be deleterious to the catalyst in catalytic reforming zone 45, and Would promote gum formation if blended directly into gasoline. A light naphtha, derived from the naphtha entering hydrofining zone 44 through lines 43 and 62, is withdrawn from hydroning zone 44 through line 63.

At least one coker distillate fraction boiling between 380 and 750 F. is passed from coker 50 to hydrocracking zone 48 through line 64, which corresponds to lines 5 and/ or 6 in FIG. 1. It is desirable that this coker distillate fraction first be hydrofined before being passed into hydrocracking zone 48, by conventional methods, to reduce the nitrogen content of the fraction to below p.p.m., and preferably to below 5 p.p.m. total nitrogen. This conveniently may be accomplished in a conventional hydrofining zone (not shown), which may operate with the same catalyst and under the same conditions as set forth below for hydrofining zone 44.

Hydrofining zone 44 may be a conventional hydrofining zone containing a conventional hydroning catalyst,

for example one wherein a coprecipitated molybdenaalumina material (e.g., such as a material prepared in accordance with the disclosures of U.S. Patent 2,432,286 to Claussen et al., or U.S. Patent 2,697,006 to Sieg) is combined With cobalt oxide, the final catalyst having a metals content equivalent to about 2% cobalt and 7% molybdenum. Representative processing conditions for removing nitrogen with this catalyst, and saturating olefins, are an LHSV of 1 to 3, 700 to 800 F., 200 to 25,000 p.s.i.g. and 1000 to 15,000 s.c.f. of hydrogen per barrel of feed stock. Ammonia and any hydrogen sulfide which may be present are removed from the effluent from hydroning zone 44 by conventional methods.

Catalytic reforming zone 45 may be a conventional catalytic reformer containing a conventional catalytic reforming catalyst such as platinum on alumina, and may operate at conventional reforming conditions.

Hydrocracking zone 48 may be a conventional hydrocracking zone containing a conventional hydrocracking catalyst, for example nickel sulfide on silica-alumina, and may operate at conventional hydrocracking conditions, for example a pressure of at least 500 p.s.i.g., preferably 800 to 3000 p.s.i.g., a temperature of from about 400 to 850 F., a hydrogen feed rate of about from 1500 to 30,000 s.c.f. per barrel, preferably from about 3000 to 15,000 s.c.f. of hydrogen per barrel of total feed, and a liquid hourly space velocity of from about 0.2 to 15, preferably from about 0.4 to 3.0. Hydrocracking zone 48 is supplied with hydrogen through line 65 from hydrogen plant 66 and through line 67 with hydrogen from catalytic reforming zone 45. Hydrogen plant 66 may be a conventional hydrogen plant supplied with fuel gas through line 68.

From hydrocracking zone 48, light ends are withdrawn through line 69, a C5 to 180 F. gasoline product is withdrawn through line 70, and a 180 to 400 F. heavy gasoline is passed through line to catalytic reforming zone 45.

Hydrogen produced in catalytic reforming zone 45 is recycled through lines 67 and 81 to hydrocracking zone 48 and hydrofining zone 44, respectively. From catalytic reforming zone 45, light ends are withdrawn through line 82, and a high octane C5-lreformate gasoline is withdrawn through line 83 as a product.

Example 2 Single- Two- Coil Coil Coker Coker Gas, LFO,l b./d 1, 958 525 Excess Butanes, b./d l, 544 1, 825 Motor Gasoline, b./d. (97.1 F-1+3 ml TEL at 10 lb. Reid Vapor Pressure) 26, 172 27, 763 Coke, Tons per Day 796 752 1 EFO (equivalent fuel oil) is amount of 10 API bunker fuel that would have equivalent heating value in Btu., assuming that one barrel of said fuel oil has a heating value of 6.3M Btu.

Although only specific modes of operation of the present process have been described, numerous variations could be made in those modes without departing from the spirit of the invention, and all such variations that fall within the scope of the appended claims are intended to be embraced thereby.

I claim:

l. In a process for converting a crude hydrocarbon feed having an end point above 950 F. to lower boiling valuable liquid products which comprises separating said l i crude hydrocarbon feed into fractionsincluding at least one naphtha fraction and at least one resduurn fraction substantially boiling above 750 F., treating said residuum fraction in a coking zone to produce coke and a liquid coker distillate, and contacting at least a portion of said liquid coker distillate in a hydrocracking zone in the presence of hydrogen and a hydrocracking catalyst under hydrocracking conditions, the improvement wherein said coking zone is a delayed coker comprising a coking furnace operated at low pressure, and at least one coke drum, and wherein said residuum is treated by separating it into a fraction boiling in the range of about 750 to 950 F. and a fraction boiling above about 950 F., passing said fraction boiling from about 750 to 950 F. to a thermal cracking furnace operated at a pressure of 300 to 1000 p.s.i.g., passing the effluent from said thermal cracking furnace through a pressure reduction zone to reduce the pressure thereon from 300 to 1000 p.s.i.g. to a low pressure below 300 p.s.i.g., passing said efuent under said reduced pressure to said coke drum, passing said fraction boiling above about 950 F. to said coking furnace, passing the efuent from said coking furnace to said coke drum, and passing to said hydrocracking zone at least a portion of the liquid Coker distillate from said coke drum having an end point below about 750 F.

2. A process as in claim 1, wherein said naphtha fraction is treated in a conventional catalytic reforming zone under conventional reforming conditions.

3. A process as in claim 1, wherein at least one gasoil fraction is separated from said crude hydrocarbon feed land is passed to said hydrocracking zone.

4. A process as in claim 1, wherein an oletinic fraction boiling in the gasoline boiling range is recovered from the eciuent from said coking zone and is substantially saturated in a conventional hydrofining zone and then is treated in a conventional reforming zone under conventional reforming conditions.

5. A process `as in claim 1, wherein said liquid coker distillate is hydroned to reduce the nitrogen content thereof to below 10 ppm. total nitrogen prior to passage thereof to said hydrocracking zone.

6. A process as in claim 2, wherein said naphtha fraction is treated in a conventional hydroning zone under conventional hydroiining conditions before being passed to said reforming zone.

References Cited by the Examiner UNlTED STATES PATENTS 2,380,897 7/45 Murphree et al 208-50 2,906,690 9/59 BIOWn 208-80 3,008,895 1l/61 HanSfOrd 20S- l l2 3,019T180 1/62 Schreiner et `al 208-80 ALPHONSO D. SULLTVAN, Primary Examiner.

Claims (1)

1. IN A PROCESS FOR CONVERTING A CRUDE HYDROCARBON FEED HAVING AN END POINT ABOVE 950*F. TO LOWER BOILING VALUABLE LIQUID PRODUCTS WHICH CONPRISES SEPARATING SAID CRUDE HYDROCARBON FEED INTO FRACTIONS INCLUDING AT LEAST ONE NAPTHA FRACTION AND AT LEAST ONE RESIDUUM FRACTION SUBSTANTIALLY BOILING ABOVE 750*F., TREATING SAID RESIDUUM FRACTION IN A COKING ZONE TO PRODUCE COKE AND A LIQUID COKER DISTILLATE, AND CONTACTING AT LEAST A PORTION OF SAID LIQUID COKER DISTILLATE IN A HYDROCRACKING ZONE IN THE PRESENCE OF HYDROGEN AND A HYDROCRACKING CATALYST UNDER HYDROCRACKING CONDITIONS, THE IMPROVEMENT WHEREIN SAID COKING ZONE IS A DELAYED COKER COMPRISING A COKING FURNACE OPERATED AT LOW PRESSURE, AND AT LEAST ONE COKE DRUM, AND WHEREIN SAID RESIDUUM IS TREATED BY SEPARATING IT INTO A FRACTION BOILING IN THE RANGE OF ABOUT 750* TO 950*F. AND A FRACTION BOILING ABOVE ABOUT 950*F., PASSING SAID FRACTION BOILING FROM ABOUT 750* TO 950*F. TO A THERMAL CRACKING FURNACE OPERATED AT A PRESSURE OF 300 TO 1000 P.S.I.G., PASSING THE EFFLUENT FROM SAID THERMAL CRACKING FURNACE THROUGH A PRESSURE REDUCTION ZONE TO REDUCE THE PRESSURE THEREON FROM 300 TO 1000 P.S.I.G. TO A LOW PRESSURE BELOW 300 P.S.I.G., PASSING SAID EFFLUENT UNDER SAID REDUCED PRESSURE TO SAID COKE DRUM, PASSING SAID FRACTION BOILING ABOVE ABOUT 950*F. TO SAID COKING FURNACE, PASSING THE EFFLUENT FROM SAID COKING FURNACE TO SAID COKE DRUM, AND PASSING TO SAID HYDROCRACKING ZONE AT LEAST A PORTION OF THE LIQUID COKER DISTILLATE FROM SAID COKE DRUM HAVING AN END POINT BELOW ABOUT 750*F.

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