CA2423946A1 - Hydrocracking process - Google Patents

Abstract

A catalytic hydrocracking process wherein a hydrocarbonaceous feedstock and a liquid recycle stream is contacted with hydrogen in a hydrocracking reaction zone at elevated temperature and pressure to obtain conversion to lower boiling hydrocarbons. The process utilizes a hydrogenation zone located in t he lower end of a hot, high pressure stripper to hydrogenate downwardly flowing liquid hydrocarbons in contact with upwardly flowing hydrogen.

Description

"HYDROCRACKING PROCESS"
BACKGROUND OF THE INVENTION
The field of art to which this invention pertains is the hydrocracking of a hydrocarbonaceous feedstock. Petroleum refiners often produce desirable products such as turbine fuel, diesel fuel and other products known as middle distillates as well as lower boiling hydrocarbonaceous liquids such as naphtha and gasoline by hydrocracking a hydrocarbon feedstock derived from crude oil, for example. Feedstocks most often subjected to hydrocracking are gas oils and heavy gas oils recovered from crude oil by distillation. A typical heavy gas oil 1o comprises a substantial portion of hydrocarbon components boiling above 371 °C, usually at least 50 percent by weight boiling above 371 °C. A typical vacuum gas oil normally has a boiling point range between 315°C and 565°C.
Hydrocracking is generally accomplished by contacting in a hydrocracking reaction vessel or zone the gas oil or other feedstock to be treated with a suitable hydrocracking catalyst under conditions of elevated temperature and pressure in the presence of hydrogen so as to yield a product containing a distribution of hydrocarbon products desired by the refiner. The operating conditions and the hydrocracking catalysts within a hydrocracking reactor influence the yield of the hydrocracked products.
2o Although a wide variety of process flow schemes, operating conditions and catalysts have been used in commercial activities, there is always a demand for new hydrocracking methods which provide lower costs and higher liquid product yields. It is generally known that enhanced product selectivity can be achieved at lower conversion per pass (60% to 90% conversion of fresh feed) through the catalytic hydrocracking zone. However, it was previously believed that any advantage of operating at below 60% conversion per pass was negligible or would .only see diminishing returns. Low conversion per pass is generally more expensive, however, the present invention greatly improves the economic benefits of a low conversion per pass process and demonstrates the unexpected advantages.
INFORMATION DISCLOSURE
US-A-5,720,872 discloses a process for hydroprocessing liquid 1o feedstocks in two or more hydroprocessing stages which are in separate reaction vessels and wherein each reaction stage contains a bed of hydroprocessing catalyst. The liquid product from the first reaction stage is sent to a low pressure stripping stage and stripped of hydrogen sulfide, ammonia and other dissolved gases. The stripped product stream is then sent to the next downstream reaction stage, the product from which is also stripped of dissolved gases and sent to the next downstream reaction stage until the last reaction stage, the liquid product of which is stripped of dissolved gases and collected or passed on for further processing. The flow of treat gas is in a direction opposite the direction in which the reaction stages are staged for the flow of liquid.
Each 2o stripping stage is a separate stage, but all stages are contained in the same stripper vessel.
International Publication No. WO 97/38066 (PCT/US 97/04270) discloses a process for reverse staging in hydroprocessing reactor systems.

BRIEF SUMMARY OF THE INVENTION
The present invention is a catalytic hydrocracking process which provides higher liquid product yields, specifically higher yields of turbine fuel and diesel oil. The present invention utilizes a hydrogenation zone located in the lower end of the hot, high pressure separator to hydrogenate the downwardly flowing hydrocarbons in contact with upwardly flowing hydrogen. Since the downwardly flowing hydrocarbons passing through the hydrogenation zone have been stripped of hydrogen sulfide and ammonia, the hydrogenation is conducted in what is known as a sweet environment which is very favorable for the removal of 1o relatively low levels of sulfur from the hydrocarbons. In addition, the upwardly flowing hydrogen effectively strips the produced hydrogen sulfide to produce ever increasingly lower sulfur hydrocarbons which are eventually removed from the bottom of the hot, high pressure separator. This resulting hydrocarbonaceous stream which is saturated with dissolved hydrogen and has 1s a very low sulfur concentration, permits the use of catalysts in the hydrocracking zone which catalysts have superior performance characteristics and are sensitive to the presence of organic and inorganic sulfur. Therefore, the process of the present invention enables the use of high performance hydrocracking catalysts which were previously unusable.
2o In one embodiment the present invention relates to a process for hydrocracking a hydrocarbonaceous feedstock which process comprises:
(a) passing a hydrocarbonaceous input stream and hydrogen to a hydrocracking zone containing hydrocracking catalyst to produce a hydrocracking effluent;
(b) combining a hydrocarbonaceous feedstock with at least one of the hydrocarbonaceous input streams or the hydrocracking effluent; (c) separating the effluent from the hydrocracking zone in a first separation zone to produce a first stream comprising hydrogen and hydrocarbons boiling at a temperature below the boiling range of the hydrocarbonaceous input stream and a s downwardly flowing liquid comprising hydrocarbonaceous compounds boiling in the range of the hydrocarbonaceous input stream; (d) simultaneously contacting the downwardly flowing liquid in the first separation zone with a hydrotreating catalyst and an upwardly flowing hydrogen stream to produce a first liquid stream comprising hydrocarbonaceous compounds boiling in the range of the 1o hydrocarbonaceous input stream; (e) recycling at least a portion of the first liquid stream to the hydrocracking zone to provide at least a portion of the hydrocarbonaceous input stream; (f) recovering a liquid hydrocarbonaceous product stream from at least a portion of the first stream contaiing hydrogen and hydrocarbons boiling at a temperature below the boiling range of the 15 hydrocarbonaceous input stream.
BRIEF DESCRIPTION OF THE DRAWINGS
Figure 1 is a simplified process flow diagram of a hydrocracking process arranged in accordance with this invention.
Figure 2 is a simplified process flow diagram of an alternate arrangement 2o for a hydrocracking process of this invention.
DETAILED DESCRIPTION OF THE INVENTION
It has been discovered that higher liquid product yields and a lower cost of production can be achieved and enjoyed in the above-described hydrocracking process. The present invention affords the more efficient and economical use of a broad range of superior hydrocracking catalysts.
The process of the present invention is particularly useful for hydrocracking a hydrocarbon oil containing hydrocarbons and/or other organic materials to produce a product containing hydrocarbons and/or other organic materials of lower average boiling point and lower average molecular weight.
The hydrocarbon feedstocks that may be subjected to hydrocracking' by the method of the invention include all mineral oils and synthetic oils (e.g., shale oil, tar sand products, etc.) and fractions thereof. Illustrative hydrocarbon 1o feedstocks include those containing components boiling above 288°-C, such as atmospheric gas oils, vacuum gas oils, deasphalted, vacuum, and atmospheric residua, hydrotreated or mildly hydrocracked residual oils, coker distillates, straight run distillates, solvent-deasphalted oils, pyrolysis-derived oils, high boiling synthetic oils, cycle oils and cat cracker distilllates. A preferred hydrocracking feedstock is a gas oil or other hydrocarbon fraction having at least 50% by weight, and most usually at least 75% by weight, of its components boiling at temperatures above the end point of the desired product, which end point, in the case of heavy gasoline, is generally in the range from 193°C to 215°
C. One of the most preferred gas oil feedstocks will contain hydrocarbon 2o components which boil above 288°C with best results being achieved with feeds containing at least 25 percent by volume of the components boiling between 315°C and 538°C.
Also included are petroleum distillates wherein at least 90 percent of the components boil in the range from 149°C to 426°C. The petroleum distillates may be treated to produce both light gasoline fractions (boiling range, for example, from 10°C to 85°C and heavy gasoline fractions (boiling range, for example, from 85°C to 204°C. The present invention is particularly suited for maximizing the yield of liquid products including middle distillate products.
In one embodiment the selected feedstock may be first introduced into a denitrification and desulfurization reaction zone together with a hot hydrocracking zone effluent at hydrotreating reaction conditions. Preferred denitrification and desulfurization reaction conditions or hydrotreating reaction conditions include a temperature from 204°C to 482°C, a pressure from 3.55 1o MPa to 17.3 MPa and a liquid hourly space velocity of the fresh hydrocarbonaceous feedstock from 0.1 hr 1 to 10 hr j with a hydrotreating catalyst or a combination of hydrotreating catalysts.
The term "hydrotreating" as used herein refers to processes wherein a hydrogen-containing treat gas is used in the presence of suitable catalysts which are primarily active for the removal of heteroatoms, such as sulfur and nitrogen and for some hydrogenation of aromatics. Suitable hydrotreating catalysts for use in the present invention are any known conventional hydrotreating catalysts and include those which are comprised of at least one Group VIII metal, preferably iron, cobalt and nickel, more preferably cobalt and/or nickel and at least one Group VI metal, preferably molybdenum and tungsten, on a high surface area support material, preferably alumina. Other suitable hydrotreating catalysts include zeolitic catalysts, as well as noble metal catalysts where the noble metal is selected from palladium and platinum. It is within the scope of the present invention that more than one type of hydrotreating catalyst be used in the same reaction vessel. The Group Vlll metal is typically present in an amount ranging from 2 to 20 weight percent, preferably from 4 to 12 weight percent. The Group VI metal will typically be present in an amount ranging from 1 to 25 weight percent, preferably from 2 to 25 weight percent. Typical hydrotreating temperatures range from 204°C to 482°C with pressures from 3.55 MPa to 17.3 MPa, preferably from 3.55 MPa to 13.9 MPa. , In another embodiment of the present invention the resulting effluent from the denitrification and desulfurization zone of the selected feedstock may be introduced into a hydrocracking zone. The hydrocracking zone may contain one or more beds of the same or different catalyst. In one embodiment, when the preferred products are middle distillates the preferred hydrocracking catalysts utilize amorphous bases or low-level zeolite bases combined with one or more Group VIII or Group VIB metal hydrogenating components. In another embodiment, when the preferred products are in the gasoline boiling range, the ~ s hydrocracking zone contains a catalyst which comprises, in general, any crystalline zeolite cracking base upon which is deposited a minor proportion of a Group VIII metal hydrogenating component. Additional hydrogenating components may be selected from Group VIB for incorporation with the zeolite base. The zeolite cracking bases are sometimes referred to in the art as 2o molecular sieves and are usually composed of silica, alumina and one or more exchangeable cations such as sodium, magnesium, calcium, rare earth metals, etc. They are further characterized by crystal pores of relatively uniform diameter between 4 and 14 Angstroms (10-1° meters). It is preferred to employ zeolites having a relatively high silica/alumina mole ratio between 3 and 12.

Suitable zeolites found in nature include, for example, mordenite, stilbite, heulandite, ferrierite, dachiardite, chabazite, erionite and faujasite.
Suitable synthetic zeolites include, for example, the B, X, Y and L crystal types, e.g., synthetic faujasite and mordenite. The preferred zeolites are those having crystal pore diameters between 8-12 Angstroms (10-'° meters), wherein the silica/alumina mole ratio is 4 to 6. A prime example of a zeolite falling in the preferred group is synthetic Y molecular sieve.
The natural occurring zeolites are normally found in a sodium form, an alkaline earth metal form, or mixed forms. The synthetic zeolites are nearly 1o always prepared first in the sodium form. In any case, for use as a cracking base it is preferred that most or all of the original zeolitic monovalent metals be ion-exchanged with a polyvalent metal and/or with an ammonium salt followed by heating to decompose the ammonium ions associated with the zeolite, leaving in their place hydrogen ions and/or exchange sites which have actually been decationized by further removal of water. Hydrogen or "decationized" Y
zeolites of this nature are more particularly described in US-A-3,130,006.
Mixed polyvalent metal-hydrogen zeolites may be prepared by ion-exchanging first with an ammonium salt; then partially back exchanging with a polyvalent metal salt and then calcining. In some cases, as in the case of 2o synthetic mordenite, the hydrogen forms can be prepared by direct acid treatment of the alkali metal zeolites. The preferred cracking bases are those which are at least 10 percent, and preferably at least 20 percent, metal-cation-deficient, based on the initial ion-exchange capacity. A specifically desirable and stable class of zeolites are those wherein at least 20 percent of the ion exchange capacity is satisfied by hydrogen ions.

The active metals employed in the preferred hydrocracking catalysts of the present invention as hydrogenation components are those of Group VIII, i.e., iron, cobalt, nickel, ruthenium, rhodium, palladium, osmium, iridium and platinum. In addition to these metals, other promoters may also be employed in conjunction therewith, including the metals of Group VIB, e.g., molybdenum and tungsten. The amount of hydrogenating metal in the catalyst can vary within wide ranges. Broadly speaking, any amount between 0.05 percent and 30 percent by weight may be used. In the case of the noble metals, it is normally preferred to use 0.05 to 2 weight percent. The preferred method for 1o incorporating the hydrogenating metal is to contact the zeolite base material with an apueous solution of a suitable compound of the desired metal wherein the metal is present in a cationic form. Following addition of the selected hydrogenating metal or metals, the resulting catalyst powder is then filtered, dried, pelleted with added lubricants, binders or the like if desired, and calcined in air at temperatures of, e.g., 371 °-648°C in order to activate the catalyst and decompose ammonium ions. Alternatively, the zeolite component may first be pelleted, followed by the addition of the hydrogenating component and activation by calcining. The foregoing catalysts may be employed in undiluted form, or the powdered zeolite catalyst may be mixed and copelleted with other relatively less 2o active catalysts, diluents or binders such as alumina, silica gel, silica-alumina cogels, activated clays and the like in proportions ranging between 5 and 90 weight percent. These diluents may be employed as such or they may contain a minor proportion of an added hydrogenating metal such as a Group VIB and/or Group VIII metal.

Additional metal promoted hydrocracking catalysts may also be utilized in the process of the present invention which comprises, for example, aluminophosphate molecular sieves, crystalline chromosilicates and other crystalline silicates. Crystalline chromosilicates are more fully described in US-A-4,363,718.
The hydrocracking of the hydrocarbonaceous feedstock or a hydrocarbonaceous input stream in contact with a hydrocracking catalyst is conducted in the presence of hydrogen and preferably at hydrocracking conditions which include a temperature from 232°C to 468°C, a pressure from 1o 3.55 MPa to 20.8 MPa, a liquid hourly space velocity (LHSV) from 0.1 to 30 hr 1, and a hydrogen circulation rate from 337 normal m3/m3 to 4200 normal m3/m3.
In accordance with the present invention, the term "substantial conversion to lower boiling products" is meant to connote the conversion of at least 10 volume percent of the fresh feedstock.
The resulting effluent from the denitrification and desulfurization reaction zone or the hydrocracking zone is transferred without intentional heat-exchange (uncooled) and is introduced into a hot, high pressure stripping zone maintained at essentially the same pressure as the denitrification and desulfurization reaction zone or hydrocracking reaction zone where it is countercurrently 2o stripped with a hydrogen-rich gaseous stream to produce a first gaseous hydrocarbonaceous stream containing hydrocarbonaceous compounds boiling at a temperature less than 371 °C, hydrogen sulfide and ammonia, and a downwardly flowing hydrocarbonaceous liquid containing hydrocarbonaceous compounds boiling at a temperature greater than 371 °C. The stripping zone is preferably maintained at a temperature in the range from 232°C to 468°C. The effluent from the denitrification and desulfurization reaction zone or hydrocracking reaction zone is not substantially cooled prior to stripping and would only be lower in temperature due to unavoidable heat loss during s transport from the reaction zone to the stripping zone. It is preferred that any cooling of the denitrification and desulfurization reaction zone effluent or the hydrocracking reaction zone prior to stripping is less than 38°C. By maintaining the pressure of the stripping zone at essentially the same pressure as the denitrification and desulfurization reaction zone or hydrocracking reaction zone is meant that any difference in pressure is due to the pressure drop required to flow the effluent stream from the reaction zone to the stripping zone. It is preferred that the pressure drop is less than .800 MPa.
The downwardly flowing hydrocarbonaceous stream containing hydrocarbonaceous compounds boiling at a temperature greater than 371 °C is 1s simultaneously contacted with a hydrotreating catalyst in a hydrogenation zone and an upwardly flowing hydrogen stream in the lower end of the hot, high pressure stripper. The hydrotreating catalyst may be selected from any known catalyst and examples of such catalysts are described hereinabove. Operating conditions in this hydrogenation zone preferably include a temperature from 20 204°C to 482°C, a pressure from 3.55 MPa to 17.3 MPa and a liquid hourly space velocity from 0.1 hr 1 to 10 hr 1. The upwardly flowing hydrogen may be supplied by a hydrogen make-up gas stream, a hydrogen-rich recycle gas stream or combinations thereof. It is preferred that the upwardly flowing hydrogen in the hydrogenation zone contains less than 50 wppm sulfur. After the downwardly flowing hydrocarbonaceous liquid passes through the hydrogenation zone, a liquid hydrocarbonaceous stream containing hydrocarbonaceous compounds boiling at a temperature greater than 371 °C, saturated with elemental hydrogen and preferably containing less than 100 s wppm sulfur is removed from the hot, high pressure stripper.
The hydrogen-rich gaseous stream is preferably supplied to the stripping zone in an amount greater than 1 weight percent of the hydrocarbonaceous feedstock. In one embodiment, the hydrogen-rich gaseous stream used as the stripping medium in the stripping zone is first introduced into a reflux heat 1 o exchange zone located in an upper end of the stripping zone to produce reflux therefor and then introducing the resulting heated hydrogen-rich gaseous stream into a lower end of the stripping zone to perform the stripping function. The hydrogen-rich gaseous stream may be introduced either below the hydrogenation zone or a portion above and below the hydrogenation zone.
15 Since the quantity of stripping gas is preferably greater than the hydrogen required in the hydrogenation zone, it is, in one embodiment, preferred that this hydrogen-rich gas is introduced in at least two locations or elevations of the hot, high pressure stripper. In the event that a high performance, low pressure drop hydrotreating catalyst is available, it is preferred in another embodiment to 2o introduce a majority, if not all, of the hydrogen-rich gas below the hydrogenation zone.
At least a portion of the first liquid hydrocarbonaceous stream containing hydrocarbonaceous compounds boiling at a temperature greater than 371 °C

recovered from the stripping zone is introduced directly into a hydrocracking zone along with added hydrogen. In a preferred embodiment, the per pass conversion in the hydrocracking zone is in the range from 15% to 45%. More preferably the per pass conversion is in the range from 20% to 40%.
The resulting first gaseous hydrocarbonaceous stream containing hydrocarbonaceous compounds boiling at a temperature less than 371 °C, hydrogen, hydrogen sulfide and ammonia from the stripping zone is preferably introduced in an all vapor phase into a post-treat hydrogenation reaction zone to hydrogenate at least a portion of the aromatic compounds in order to improve 1o the quality of the middle distillate, particularly the jet fuel. The post-treat hydrogenation reaction zone may be conducted in a downflow, upflow or radial flow mode of operation and may utilize any known hydrogenation catalyst. The effluent from the post-treat hydrogenation reaction zone is preferably cooled to a temperature in the range from 4°C to 60°C and at least partially condensed to ~5 produce a second liquid hydrocarbonaceous stream which is recovered and fractionated to produce desired hydrocarbon product streams and to produce a second hydrogen-rich gaseous stream which is bifurcated to provide at least a portion of the added hydrogen introduced into the hydrocracking zone as hereinabove described and at least a portion of the first hydrogen-rich gaseous 2o stream introduced in the stripping zone. Fresh make-up hydrogen may be introduced into the process at any suitable and convenient location but is preferably introduced into the stripping zone and more preferably below the hydrogenation zone in the stripping zone. Before the second hydrogen-rich gaseous stream is introduced into the hydrocracking zone, it is preferred that at least a significant portion, at least 90 weight percent, for example, of the hydrogen sulfide is removed and recovered by means of known, conventional methods. In a preferred embodiment, the hydrogen-rich gaseous stream introduced into the hydrocracking zone contains less than 50 wppm hydrogen sulfide.
DETAILED DESCRIPTION OF THE DRAWINGS
With reference now to Figure 1, a feed stream comprising vacuum gas oil and heavy coker gas oil is introduced into the process via line 1 and admixed with a hereinafter-described effluent from hydrocracking zone 31 transported via line 32. The resulting admixture is transported via line 2 into hydrotreating zone 3. The resulting effluent from hydrotreating zone 3 is transported via line 4 and introduced into stripping zone 5. A vaporous stream containing hydrocarbons and hydrogen passes upward in stripping zone 5 and contacts heat-exchanger 25 and at least a portion thereof is removed from stripping zone 5 via line 7 and 1 s introduced into post-treat hydrotreating zone 8. A downwardly flowing hydrocarbonaceous liquid enters and passes through hydrotreating zone 40. A
liquid hydrocarbonaceous stream is removed from stripping zone 5 via line 6 and is introduced into hydrocracking zone 31 via line 6 and line 30. A gaseous effluent stream is removed from post-treat hydrotreating zone 8 via line 9 and is 2o introduced into heat-exchanger 10. The resulting cooled effluent from heat-exchanger 10 is transported via line 11 and introduced into vapor-liquid separator 12. A hydrogen-rich gaseous stream containing acid gas compounds is removed from vapor-liquid separator 12 via line 17 and is introduced into acid gas recovery zone 18. A lean solvent is introduced via line 35 into acid gas recovery zone 18 and contacts the hydrogen-rich gaseous stream in order to dissolve an acid gas. A rich solvent containing acid gas is removed from acid gas recovery zone 18 via line 36 and recovered. A hydrogen-rich gaseous stream containing a reduced concentration of acid gas is removed from acid gas recovery zone 18 via line 19 and is admixed with fresh make-up hydrogen which is introduced via line 20. The resulting admixture is transported via line 21 and is introduced into compressor 22. A resulting compressed hydrogen-rich gaseous stream is transported via line 23 and at least a portion is recycled via line 29 and line 30 to hydrocracking zone 31. Another portion of the hydrogen-1o rich gaseous stream is transported via line 24 and is introduced into heat-exchanger 25. The resulting heated hydrogen-rich gaseous stream is removed from heat-exchanger 25 via line 26 and is introduced into heat-exchanger 27.
The resulting heated hydrogen-rich gaseous stream is removed from heat-exchanger 27 and transported via line 28 and introduced into stripping zone 5 at a location below hydrotreating zone 40. An aqueous stream is introduced via line 33 and contacts the flowing stream in line 9 and is subsequently introduced into vapor-liquid separator 12 as hereinabove described. An aqueous stream containing water-soluble salts is removed from vapor-liquid separator 12 via line 34 and recovered. A liquid stream containing hydrocarbonaceous compounds is 2o removed from vapor-liquid separator 12 via fine 13, reduced in pressure and introduced into separation zone 14. A gaseous stream containing hydrogen and normally gaseous hydrocarbons is removed from separation zone 14 via line 15.
A liquid stream containing hydrocarbons is removed from separation zone 14 via line 16 and recovered.

With reference now to Figure 2, a feed stream comprising vacuum gas oil and heavy coker gas oil is introduced into the process via line 51 and admixed with a hereinafter-described recycle stream transported via line 56. The resulting admixture is transported via line 56 and is contacted with a hydrogen-rich gaseous stream provided via line 129 and the resulting admixture is introduced via line 130 into hydrotreating zone 131 which does not significantly alter or hydrocrack the feed stream, but converts heteroatom compounds to produce ammonia and hydrogen sulfide. The resulting effluent from hydrotreating zone 131 is transferred via line 52 and introduced into 1 o hydrocracking zone 53. The resulting effluent from hydrocracking zone 53 is transported via line 54 and introduced into stripping zone 58. A vaporous stream containing hydrocarbons and hydrogen passes upward in stripping zone 58 and contacts packing 57 and then heat-exchanger 125 and at least a portion thereof is introduced into post-treat hydrotreating zone 132. A downwardly flowing hydrocarbonaceous liquid enters and passes through hydrotreating zone 55. A liquid hydrocarbonaceous stream is removed from stripping zone 58 via line 56 and is introduced into hydrotreating zone 131 via line 56 and line 130.
Another liquid hydrocarbonaceous stream is removed from stripping zone 58 via lirie 56 and line 137 and introduced into separation zone 114. A gaseous 2o effluent stream is removed from ~ post-treat hydrotreating zone 132 via line 59 and is introduced into heat-exchanger 110. The resulting cooled effluent from heat-exchanger 110 is transported via line 111 and introduced into vapor-liquid separator 112. A hydrogen-rich gaseous stream containing acid gas compounds is removed from vapor-liquid separator 112 via line 117 and is introduced into acid gas recovery zone 118. A lean solvent is introduced via line 135 into acid gas recovery zone 118 and contacts the hydrogen-rich gaseous stream in order to dissolve an acid gas. A rich solvent containing acid gas is removed from acid gas recovery zone 118 via line 136 and recovered. A
hydrogen-rich gaseous stream containing a reduced concentration of acid gas is removed from acid gas recovery zone 118 via line 119 and is admixed with fresh make-up hydrogen which is introduced via line 120. The resulting admixture is transported via line 121 and is introduced into compressor 122. A resulting compressed hydrogen-rich gaseous stream is transported via line 123 and at least a portion is recycled via line 129 and line 130 to hydrotreating zone 131.
1o Another portion of~the hydrogen-rich gaseous stream is transported via line and is introduced into heat-exchanger 125. The resulting heated hydrogen-rich gaseous stream is removed from heat-exchanger 125 via line 126 and at least a portion is introduced into stripping zone 58 at a location below packing 57 via line 128 and at least another portion is transported via line 127 and introduced into stripping zone 58 at a location below hydrotreating zone 55. An aqueous stream is introduced via line 133 and contacts the flowing stream in line 59 and is subsequently introduced into vapor-liquid separator 112 as hereinabove described. An aqueous stream containing water-soluble salts is removed from vapor-liquid separator 112 via line 134 and recovered. A liquid stream 2o containing hydrocarbonaceous compounds is removed from vapor-liquid separator 112 via line 113, reduced in pressure and introduced into separation zone 114. A gaseous stream containing hydrogen and normally gaseous hydrocarbons is removed from separation zone 114 via line 115. A liquid stream containing hydrocarbons is removed from separation zone 114 via line 116 and recovered.

A portion of a hydrocracker feedstock having the characteristics presented in Table 1 is hydrocracked in a conventional single stage hydracracker at operating conditions presented in Table 2 to yield the products described in Table 3. Another portion of the same hydrocracker feedstock is hydrocracked in a hydrocracker of the present invention using the same type of catalyst as the base case at operating conditions presented in Table 2 to yield 1o the products described in Table 3. Yields are calculated based on fresh feed at start of run conditions.

80/20 Blend Stra~ht Run Vacuum Gas Oil-Coker Gas Oil Specific Gravity, 15°C 0.928 Distillation, Volume Percent IBP, -C 351 Sulfur, weight percent 3.01 Nitrogen, PPM 1256 Bromine Number 7.5 Heptane Insolubles, weight <0.05 percent Conradson Carbon, weight 0.36 percent Nickel and Vanadium, PPM 0.4 Low Conversion Per Pass with Improved Flowscheme Base Case Yields Reactor Operating Conditions High Pressure Separator Pressure,16 11.8 Mpa Liquid Hourly Space Velocity Hydrotreating Zone 2.18 1.13 Hydrocracking Zone 0.93 3.0 Overall 0.65 0.82 Combined Feed Ratio **1.5 ***3.0 H2/Fresh Feed, m3/ m3 1955 1955 Conversion, Per Pass*, % 60 30 Total (Gross) Conversion, %* 100 100 Number of Gas Quench Points 3 0 Maximum Reactor ~T,C HT/HC 30/18 36/30 * Conversion to 382°C End Point Distillate and Lighter ** Recycle Liquid to HT first then to HC
*** Recycle Liquid to HC first then to HT

Base Case Invention Wt. % Vol. % Wt. % Vol.

NH3 0.15 0.15 H2S 3.20 3.20 C~-C4 3.68 2.97 Light Naphtha (C5-C6) 6.32 8.76 5.08 7.04 Heavy Naphtha (C~-127C)10.38 12.87 7.68 9.52 Kerosine (127C-288C) 50.16 58.15 48.34 55.92 Diesel (288C-382C) 28.72 31.98 35.11 39.09 Total Middle Distillate78.88 90.13 83.45 95.01 C5+ Total 95.58 111.76 96.21 111.57 C4+ Total 98.20 116.01 98.32 115.00 Chemical H2 Consumption2.61 284 2.53 276 m3/ m3 From the above tables it is apparent that the present invention is able to operate at a pressure of 11.8 MPa or. approximately one fourth less than the base case, utilizes a hydrocracking reactor having 30% less internal volume as well as 20% less catalyst inventory. Because of the lower hydrocracking reactor zone operating severity in the present invention, the conversion per pass is reduced from 60% to 30%. These enumerated changes used in the present invention provide a lower cost hydrocracking process as well as providing an increased yield of total middle distillate product. The present invention also has a 8 m3/m3 lower chemical hydrogen consumption and a 50% less hydrogen loss to fuel gas.
The foregoing description, drawings and examples clearly illustrate the advantages encompassed by the process of the present invention and the benefits to be afforded with the use thereof.

Claims (8)

CLAIMS:

1. A process for hydrocracking a hydrocarbonaceous feedstock which process comprises:
(a) passing a hydrocarbonaceous input stream and hydrogen to a hydrocracking zone containing hydrocracking catalyst to produce a hydrocracking effluent;
(b) combining a hydrocarbonaceous feedstock with at least one of the hydrocarbonaceous input streams or the hydrocracking effluent;
(c) separating the effluent from the hydrocracking zone in a first separation zone to produce a first stream comprising hydrogen and hydrocarbons boiling at a temperature below the boiling range of the hydrocarbonaceous input stream and a downwardly flowing liquid comprising hydrocarbonaceous compounds boiling in the range of the hydrocarbonaceous input stream;
(d) simultaneously contacting the downwardly flowing liquid in the first separation zone with a hydrotreating catalyst and an upwardly flowing hydrogen stream to produce a first liquid stream comprising hydrocarbonaceous compounds boiling in the range of the hydrocarbonaceous input stream;
(e) recycling at least a portion of the first liquid stream to the hydrocracking zone to provide at least a portion of the hydrocarbonaceous input stream;
(f) recovering a liquid hydrocarbonaceous product stream from at least a portion of the first stream contaiing hydrogen and hydrocarbons boiling at a temperature below the boiling range of the hydrocarbonaceous input stream.

2. The process of Claim 1 wherein prior to separation in the first separation zone the effluent from the hydrocracking zone and the hydrocarbonaceous feedstock pass to a denitrification and desulurization reaction zone containing a catalyst and the denitrification and desulfurization zone effluent undergoes separation to produce the first stream and the downwardly flowing liquid.

3. The process of Claims 1 and 2 wherein the denitrification and desulfurization reaction zone effluent or the hydrocracking effluent passes directly to the first separation zone which comprises a hot, high pressure stripper utilizing a hot hydrogen-rich stripping gas to produce the first stream comprising hydrogen and hydrocarbonaceous compounds boiling at a temperature below the boiling range of the hydrocarbonaceous feedstock, and to produce the downwardly flowing liquid comprising hydrocarbonaceous compounds boiling in the range of the hydrocarbonaceous feedstock.

4. The process of Claim 3 wherein the effluent from the hydrocracking zone passes to the denitrification and desulfurization zone and at least a portion of the downwardly flowing liquid passes to the hydrocracking zone as the hydrocarbonaceous input stream.

5. The process of Claim 3 wherein the first stream comprising hydrogen and hydrocarbonaceous compounds boiling at a temperature below the boiling range of the hydrocarbonaceous feedstock passes to an aromatic saturation zone containing hydrogenation catalyst.

6. The process of Claims 1-5 wherein the hydrocarbonaceous feedstock boils in the range from 315°C to 538°C.

7. The process of Claims 3-6 wherein the hot, high pressure stripper is operated at a temperature no less than 38°C below the outlet temperature of the denitrification and desulfurization reaction zone and at a pressure no less than 689 kPa below the outlet pressure of the denitrification and desulfurization reaction zone.

8. The process of Claims 1-7 wherein the hydrocracking zone is operated at a conversion per pass in the range from 15 to 45.