Classifications

• • C—CHEMISTRY; METALLURGY
  • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
  • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
  • C10G47/00—Cracking of hydrocarbon oils, in the presence of hydrogen or hydrogen- generating compounds, to obtain lower boiling fractions
• • C—CHEMISTRY; METALLURGY
  • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
  • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
  • C10G67/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only

Definitions

• ABSTRACT Residue-containing petroleum oils are converted into lighter [52] U.S.Cl ..208/61,208/ 164 products y a combination of catalytic hydrocracking and [51 Int. Cl. ..Cl0g 13/00, ClOg 37/02 catalytic cracking [58] Field of Search ..208/59, 61, 164
• This invention relates to the catalytic treatment of heavy hydrocarbon materials and more particularly to a process which produces substantially complete conversion of said heavy hydrocarbon materials to lower boiling hydrocarbons and selectivity in such conversion to lower boiling hydrocarbons which boil within a particularly preferred boiling range.
• the still residue is a heavy hydrocarbon oil rich in tar and asphalt and having a relatively high concentration of metals.
• Attempts to convert still residues such as a vacuum residuum into lighter materials by means of catalytic processes have not been particularly successful as the tar and asphalt deposit on the catalyst producing a coke layer on the catalyst preventing contact of the catalyst and oil.
• the metals will deposit on the catalyst causing its deactivation.
• the most popular method for converting residua to lighter materials is coking, in which process the oil is heated and retained at elevated temperature until a substantial portion thereof is converted to coke and the balance to a lighter liquid.
• disposal of the coke so formed can present a problem.
• much of the still residue produced in petroleum refineries is sold as residual fuel" but even this is no longer a suitable use because of its high sulfur content.
• our invention provides a process for the conversion of a residuecontaining petroleum fraction into lighter products which comprises maintaining in a hydrocracking zone a first catalytic zone below and a second catalytic zone above a point of entry into said hydrocracking zone, introducing the residuecontaining petroleum fraction through said point of entry into said hydrocracking zone at a temperature between about 600 and 850 F.
• the process of our invention may be used for the treatment of residue-containing fractions such as atmospheric residua,
• vacuum residua vacuum residua, visbreaker bottoms, whole crude such as San Ardo Crude, shale oil, tar sand oil and the like.
• a split flow hydrocracking process which comprises introducing a heavy hydrocarbon charge stock in downward flow into a hydrocracking catalyst zone, said catalyst zone comprising a first catalyst zone below and a second catalyst zone above the point of entry of the heavy hydrocarbon charge stock, introducing hydrogen into said first catalyst zone in countercurrent relationship to the flow of said heavy hydrocarbon charge stock, maintaining a lower boiling liquid in the second catalyst zone, recovering a high boiling effluent from the first catalyst zone and recovering lower boiling hydrocarbons from the second catalyst zone.
• the heavy hydrocarbon charge stock is introduced into a hydrocracking catalyst zone herein defined to include a first catalyst zone in downflow relationship to the downward flow of the heavy hydrocarbon charge stock, a second catalyst zone above the point of entry of the heavy hydrocarbon charge stock and in upfiow relationship to the lower boiling hydrocarbons which proceed from the first catalyst zone into the second catalyst zone.
• a hydrocracking catalyst zone herein defined to include a first catalyst zone in downflow relationship to the downward flow of the heavy hydrocarbon charge stock, a second catalyst zone above the point of entry of the heavy hydrocarbon charge stock and in upfiow relationship to the lower boiling hydrocarbons which proceed from the first catalyst zone into the second catalyst zone.
• the second catalyst zone is in upfiow relationship to the flow of the hydrogen containing gas and in upfiow relationship to the volatile hydrocarbon and entrained liquid hydrocarbon which proceed from the first catalyst zone into a second catalyst zone.
• the word above is used to define a flow relationship with the first catalyst zone, which relationship provides for the flow of hydrogen, volatile hydrocarbons and entrained lower boiling liquid hydrocarbons from the first catalyst zone in countercurrent relationship with the downward flow of the heavy hydrocarbon charge stock into a second catalyst zone.
• the second catalyst zone can be located directly in a space dimension above the first catalyst zone such as when the first and second catalyst zone are present in a vertical reactor with an intermediate point of entry for the heavy hydrocarbon charge stock.
• the second catalyst zone can be present as a separate reactor which is connected to the first reactor by conduit means, although it is preferred in carrying out the process of this invention to use a vertical reactor wherein the first catalyst zone and second catalyst zone are present in the same reactor.
• a catalyst which has hydrocracking activity under process conditions of temperature, pressure and space velocity which are utilized during the process.
• the catalyst in the first catalyst zone can be either. the same or different than the catalyst-present in the second catalyst zone.
• the heavy hydrocarbon charge stock upon entry to the catalyst zone proceeds downwardly in downflow relationship tothe first catalyst zone.
• Hydrogen is introduced into the first catalyst zone at or near the lower extremity and/or at intermediate points in said first catalyst zone in countercurrent relationship to the hydrocarbon flow through the first catalyst zone and in upfiow relationship to the second catalyst zone, the volatile hydrocarbons and the lower boiling liquid hydrocarbons hereinafler referred to as liquid proceed into the second catalyst zone.
• the volatile hydrocarbons and the liquids which are present in the second catalyst zone proceed from the second catalyst zone and are recovered by conventional means such as by cooling of the hydrocarbon vapors and liquid.
• the hydrogen which proceeds from the second catalyst zone can then be recycled together with fresh hydrogen in the first catalyst zone.
• hydrogen optionally can be blended with the heavy hydrocarbon charge stock and introduced at ambient temperature or higher such as temperatures up to hydrocracking temperatures into the catalyst zone.
• liquid is maintained in the second catalyst zone.
• a liquid is maintained in the second catalyst zone by the rate of introduction of hydrogen into the first catalyst zone by any of the means set forth above for the introduction of hydrogen.
• hydrogen gas rates of at least 3,000 SCF per barrel of charge preferably from 3,000 SCF per barrel up to about 25,000 SCF per barrel are required in the first catalyst zone.
• the hydrogen need not be pure and gases containing more than about 65 volume percent hydrogen may be used.
• the term hydrogen is also intended to include dilute hydrogen, reformer by-product hydrogen, hydrogen produced by the partial oxidation of hydrocarbon materials followed by shift conversion and electrolytic hydrogen.
• the hold-up of the liquid hydrocarbon charge stock in the first catalyst zone can be varied somewhat by varying the upward flow of hydrogen. In general it is preferred to have high liquid hold-up, that is a hold-up of hydrocarbon charge stock which provides for maximum catalytic effectiveness for the conversion of the charge stock to lower boiling hydrocarbons.
• the lower boiling liquid which is maintained in the second catalyst zone in general is derived from the heavy hydrocarbon material, and in general is a lower boiling hydrocarbon which is present initially in the heavy hydrocarbon charge stock and/or which is formed during the process.
• the liquid material has a boiling point below 850 F. It is preferred that the liquid which is present in the second catalyst zone have at least 90 percent by weight of the liquid boiling below 850 F. more preferably at least about 97 percent by weight boiling below 850 F.
• the first stage of the process of this invention is utilized for the hydrocracking of heavy hydrocarbon charge stocks which term hydrocracking is herein defined to mean destructive hydrogenation in which a substantial portion of the product boils at a temperature below the initial boiling point of the charge heavy hydrocarbon material.
• hydrocracking is herein defined to mean destructive hydrogenation in which a substantial portion of the product boils at a temperature below the initial boiling point of the charge heavy hydrocarbon material.
• percent conversions by weight per single pass of the 850 F.+ material of the charge stock varies from about to 80 percent more preferably from about to 60 percent.
• the hydrocracking conditions as to pressure, temperature and space velocity can be varied over a wide range, the conditions utilized being those which in combination produce substantial conversion of the heavy hydrocarbon charge stock to lower boiling hydrocarbons.
• the first and second catalyst zone conditions that are utilized in the split flow process of this invention are in general temperatures of from about 600 F. to about 850 F., preferably 725 to 840 F.; pressure of from about 500 to about 5,000 psig, preferably 1,500 to 2,000 psig and liquid hourly space velocities of from about 0.05 to about 10, preferably 0.25 to 2.5, volumes of feed per volume of catalyst per hour.
• the gas rates in the first and second catalyst zones will differ depending upon the amount of hydrogen which is blended together with the heavy hydrocarbon charge stock prior to the introduction into the catalyst zone and/or hydrogen consumed in the process.
• hydrogen gas rates in the second catalyst zone may be different than the hydrogen rates in the first catalyst zone.
• liquid hourly space velocity in the second catalyst zone will be greater than that in the first catalyst zone.
• temperature and pressure can be different.
• the hydrocracking catalyst utilized for the conversion of the aforementioned hydrocarbon charge stocks can be crystalline metallic alumino-silicate zeolite, having a platinum group metal (e.g. platinum or palladium) or an iron group metal alone or in conjunction with a Group VI metal, their compounds and mixtures thereof eg cobalt oxide and molyb- 1 denum oxide or nickel sulfide and tungsten sulfide deposited thereon or composited therewith.
• platinum group metal e.g. platinum or palladium
• an iron group metal alone or in conjunction with a Group VI metal
• crystalline zeolites are characterized by their highly ordered crystalline structure and uniformly dimensioned pores, and have an alumino-silicate anionic cage structure wherein alumina and silica tetrahedra are intimately connected to each other so as to provide a large number of active sites, with the uniform pore openings facilitating entry of certain molecular structures. It has been found that crystalline alumino-silicate zeolites, having effective pore diameter of about 6 to 15, preferably 8 to 15 angstrom units, when composited with the platinum group metal, and particularly after base exchange to reduce the alkali metal oxide (e.g. Nat-,0) content of thezeolite to less than about 10 wt. preferably less than 2.0%, are effective hydrocracking catalysts.
• the support will also contain at least one amorphous inorganic oxide such as silica,
• Such composite supports preferably contain about 15-45 percent zeolite.
• the catalyst support may be totally amorphous inorganic oxide.
• Suitable such carriers or supports include acidic supports such as: silica-alumina, silica-magnesia, and other well-known cracking catalyst bases; the acidic clays; fluorided alumina; and mixtures of inorganic oxides, such as alumina, silica, zirconia, and titania, having sufficient acidic properties providing high cracking activity.
• each catalyst zone contains an amorphous support and the catalyst in the second zone contains a crystalline zeolite of low alkali metal content in the support.
• Hydrogen is separated from the effluent from the second catalytic zone and if desired may be recycled'to the first catalytic zone with or without purification for the removal of compounds such as hydrogen sulfide and/or ammonia.
• Lower boiling hydrocarbons eg those boiling up to about 525-550 F. are also removed from the second catalytic zone effluent and the balance is subjected to catalytic cracking as is the effluent from the first catalytic stage. Since the effluent from the second catalytic zone or overhead is high in saturates and the effluent from the first catalytic zone or bottoms is high in aromatics, they are subjected to different conditions of catalytic cracking.
• a zeolite cracking catalyst in a fluid catalytic cracking unit comprising a reactor, a regenerator and at least two elongated reaction zones or risers where the reactor contains a dense phase and a dilute phase of the catalyst.
• an FCCU with two risers is employed with the operating conditions in the risers including a temperature of 800-l,l50 F., conversion of 30-80 volume percent and space velocities in the overhead riser and the bottoms riser being 10-100 w/hr/w and 50-200 w/hr/w, respectively.
• the cracking of the overhead and the bottoms is restricted to the risers by discharging the efiluent from both risers into the dilute phase of catalyst in the reactor vessel.
• the reactor vessel in this case is utilized as a disengaging space with substantially no cracking taking place therein.
• the overhead is subjected to both riser and dense phase cracking while the cracking of the bottoms is limited to its riser.
• the effluent from the bottoms riser is discharged into the dilute phase of catalyst, the effluent from the overhead riser is discharged into the dense phase of catalyst and the vaporous reaction mixture from the overhead riser is passed through the dense phase of catalyst under catalytic cracking conditions effecting an additional conversion of 5-30 volume percent with the total per pass conversion of the overhead not exceeding volume percent.
• the conversion in the overhead riser may be lower, equal to or higher than that in the bottoms riser.
• the overhead is subjected only to riser cracking while the bottoms is cracked in both the riser and the dense phase of catalyst.
• the effluent from the overhead riser is discharged directly into the dilute phase of catalyst in the reactor vessel, while the effluent from the bottoms riser is discharged into the dense phase of catalyst and passed through this dense phase under catalytic cracking conditions effecting an additional conversion of 5-30 volume percent.
• the per pass conversion of the bottoms does not exceed 80 volume percent.
• both the overhead and the bottoms are subjected to both riser cracking and dense phase bed cracking by discharging the effluent from both risers into the dense phase of catalyst and passing them therethrough under catalytic cracking conditions to effect an additional conversion of 5-30 percent.
• the total conversion of all oils passing through the catalytic cracking unit does not exceed 80 volume percent.
• the overhead is subjected only to riser cracking and the bottoms is subjected to both riser and dense phase cracking.
• a virgin gas oil may be introduced into the riser with the overhead and the unconverted oil which ordinarily is recycled to the cracking unit is introduced into a separate riser with the bottoms.
• a crude oil is fractionated at atmospheric pressure to produce naphtha, kerosene, atmospheric gas oils and an atmospheric residuum
• the atmospheric residuum is subjected to split flow hydrocracking
• the atmospheric gas oil is subjected to catalytic cracking
• the overhead from the split flow hydrocracking is combined with the atmospheric gas oil as fresh feed to the catalytic cracking zone
• the unconverted feed is combined with the bottoms from the split flow hydrocracking and catalytically cracked under conditions such that the conversion of the lighter material is at least as great as that of the heavier material and may be as much as 30 percent more.
• the catalyst employed in the instant invention comprises a large pore crystalline aluminosilicate customarily referred to as a zeolite and an active metal oxide, as exemplified by silicaalumina gel or clay.
• the zeolites employed as cracking catalysts herein possess ordered rigid three-dimensional structures having uniform pore diameters within the range of from about 5 to about 15 A.
• the crystalline zeolitic catalysts employed herein comprise about 1 to 25 wt. zeolite, about to 50 wt. alumina and the remainder silica.
• zeolites are those known as zeolite X and zeolite Y wherein at least a substantial portion of the original alkali metal ions have been replaced with such cations as hydrogen and/or metal or combination of metals such as barium, calcium, magnesium, manganese or such rare earth metals, for example. cerium, lanthanum, neodymium, praseodymium, samarium and yttrium.
• the overhead and bottoms are introduced into elongated reaction zones which are operated to effect a lower conversion of the bottoms stream.
• a two riser FCCU is employed.
• the operating conditions for both the overhead riser and the bottoms riser include an operating temperature of 800-l,150 F., preferably 840l, 000 F. and a conversion per pass of 30-80 percent, preferably 40-75 percent.
• Other operating conditions within the risers include a residence time of 2-20 seconds, preferably 3-10 seconds and a vapor velocity of -50 ft/sec, preferably -40 ft/sec.
• the space velocity in the overhead riser is 10-100 w/hr/w, preferably 40-90 w/hr/w and the space velocity in the bottoms riser is 50-200 w/hr/w, preferably 75-150 w/hr/w.
• the conversion per pass in the bottoms riser is 0-30 percent lower than the conversion in the overhead riser with the overall conversion in the overhead riser not exceeding 80 volume percent.
• the operating conditions within the dense phase include a temperature of 800-l,150 F., a vapor velocity of 0.5-4 ft/sec. preferably 1.3-2.2 ft/sec and a space velocity of l-40 w/hr/w, preferably 3-25 w/hr/w.
• the vaporout reaction products from a riser which passes through the dense phase of catalyst obtains a further conversion of 5-30 volume percent.
• Another feature of our invention is that by introducing the bottoms product into the catalytic cracking unit, considerably more carbon than usual is introduced into the catalyst bed thereby permitting greater deposition of carbon on the catalyst which in turn permits operation with a regenerated catalyst having a carbon level up to about 4.0 weight percent.
• a regenerated catalyst having a carbon level up to about 4.0 weight percent.
• Another feature of our process is that the more easily cracked feed to the cracking stage can be introduced and reacted separately from the more difficultly cracked material under less severe conditions thereby avoiding overcracking with the undesirable production ofgases obtained in conventional processes where the feed is a single mixture of several streams or where the only difference between two or more feeds lies in the boiling range and not in the type of hydrocarbons in the feeds.
• EXAMPLE I In this example, a South Louisiana reduced crude having an API Gravity of 2 l 1, a sulfur content of 0.48 wt. and a Conradson Carbon Residue of4. 14 wt.% is hydrocracked by being passed downwardly with hydrogen through a bed of pelleted catalyst containing 3.4 wt. N10 and 15.9 wt. Mo 0;, supported on alumina. Reaction conditions and yield data for three runs are tabulated below:
• EXAMPLE 11 In this example the countercurrent split flow technique of the first stage of the process of our invention is employed using the same charge and catalyst as in Example 1. Reaction conditions and yield data for three runs are tabulated below:
• SCFB 234 314 552 Overhead Product Gravity. AP1 34.0 35.2 36.3 Sulfur. Wt. 0.040 0.004 0.003 Overhead Product Dist. F. Vol.
• Example 11 shows the superiority of counter-current split flow hydrocracking over conventional downfiow hydrocracking.
• EXAMPLE III In this example the products from Runs 3 and 6 are catalytically cracked under substantially identical conditions using a cracking catalyst containing 2.0 wt. cerium, 0.93 wt. lanthanum. 18.0 wt. 7: decationized zeolite Y, 0.94 wt. sodium. 34.1 wt. alumina and the balance silica and having a surface area of 329 m /g and a pore volume of 0.72 cc/g. ln Runs 7. 8 and 9 the feeds are the product from Run 3, the overhead from Run 6. and the bottoms from Run 6 respectively. The extent of cracking is shown by the gas chromatographic analysis of the charges and the products appearing in Table 3.
• EXAMPLE IV In this example the overhead and bottoms from Run 6 are introduced separately into a dual riser fluid catalytic cracking unit in which the fresh feed is a gas oil blend and in which 430 F.+ product is recycled.
• the cracking catalyst is the same as that used in Example 111.
• the proportion of feed to the risers is given below as percent of total feed to the unit.
• Riser No. 1 is operated at a temperature of 860 F. and Riser No. 2 at a temperature of 940 F. with substantially all of the cracking taking place in the risers.
• the conversion of 430 F material into 430 F. material in Riser l is 67.0 volume percent and in Riser 2 49.0 percent giving an overall conversion of of 58.0 percent.
• the total yield of debutanized naphtha is 49.2 volume percent basis feed having a Research Octane No. (with 3 cc TEL/gal) of96.5.
• a process for the conversion of a residue-containing petroleum fraction into lighter products which comprises maintaining in a hydrocracking zone a first catalytic zone below and a second catalytic zone above a point of entry into said hydrocracking zone, introducing the residue-containing petroleum fraction through said point of entry into said hydrocracking zone at a temperature between about 600 and 850 F.

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• Oil, Petroleum & Natural Gas (AREA)
• Engineering & Computer Science (AREA)
• Chemical Kinetics & Catalysis (AREA)
• General Chemical & Material Sciences (AREA)
• Organic Chemistry (AREA)
• Production Of Liquid Hydrocarbon Mixture For Refining Petroleum (AREA)
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• 1970
  • 1970-12-24 US US101444A patent/US3671420A/en not_active Expired - Lifetime
• 1971
  • 1971-07-08 CA CA117,778A patent/CA960982A/en not_active Expired
  • 1971-10-02 DE DE2149370A patent/DE2149370C3/de not_active Expired
  • 1971-11-15 ZA ZA717670A patent/ZA717670B/xx unknown
  • 1971-11-16 GB GB5307171A patent/GB1346336A/en not_active Expired
  • 1971-11-26 ES ES397401A patent/ES397401A1/es not_active Expired
  • 1971-12-15 FR FR7145004A patent/FR2118919B1/fr not_active Expired
  • 1971-12-16 BE BE776823A patent/BE776823A/xx unknown
  • 1971-12-20 SE SE7116358A patent/SE372035B/xx unknown
  • 1971-12-20 BR BR8434/71A patent/BR7108434D0/pt unknown
  • 1971-12-22 IT IT32736/71A patent/IT944220B/it active
  • 1971-12-23 NL NL7117718A patent/NL7117718A/xx unknown

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