US2710827A - Fluid hydroforming process - Google Patents

Fluid hydroforming process Download PDF

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US2710827A
US2710827A US279723A US27972352A US2710827A US 2710827 A US2710827 A US 2710827A US 279723 A US279723 A US 279723A US 27972352 A US27972352 A US 27972352A US 2710827 A US2710827 A US 2710827A
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catalyst particles
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Edward J Gornowski
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ExxonMobil Research and Engineering Co
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    • CCHEMISTRY; METALLURGY
    • C10PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10GCRACKING 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
    • C10G35/00Reforming naphtha
    • C10G35/04Catalytic reforming
    • C10G35/10Catalytic reforming with moving catalysts
    • C10G35/14Catalytic reforming with moving catalysts according to the "fluidised-bed" technique

Description

June 14, 1955 Filed April 1, 1952 FLUID HYDROFORMING PROCESS 3 Sheets-Sheet 1 1 plODUCJ" FLUE GAs E2 u l '16 H;
HEATUL Qadvcui. GIL
FUEL GAS Edward J Garages-5 54.6 firm/ember MW Qbbofneg United States Patent 2,710,827 Patented June 14, 1955 FLUID HYDROFORMJNG PROCESS Application April 1, 1952, Serial No. 279,723
6 Claims. (Cl. 196-49) This invention relates to the catalytic conversion of hydrocarbon fractions boiling within the motor fuel boiling range of low knock rating into high octane number motor fuels rich in aromatics and particularly to a process whereby such a conversion is effected by the fluidized solids technique.
It is well known that petroleum naphthas or hydrocarbon fractions boiling within the motor fuel range can be subjected to reforming operations to yield liquid products boiling within the gasoline range and possessing improved octane numbers. Reforming operations employing catalysts, especially hydroforming and aromatization processes are widely used in the petroleum industry. By hydroforming is ordinarily meant a process wherein naphtha fractions are treated at elevated temperatures and pressures in the presence of certain solid catalysts and hydrogen whereby the hydrocarbon fraction is increased in aromaticity with no net consumption of hydrogen. The term aromatization when used broadly refers to conversions which increase the aromaticity of the hydrocarbon fraction treated. As generally used in the petroleum industry, aromatization is a process in which hydrocarbon fractions are treated at elevated tem peratures in the presence of solid catalyst particles and in the presence or absence of added hydrogen, usually at pressures lower than those employed in hydroforming, for the purpose of increasing the aromaticity of the hydrocarbon fraction.
Catalytic reforming processes are usually carried out at temperatures of about 7504150 F. in the pressure range of about -3000 lbs. per sq. inch and in the presence of such catalysts as molybdenum oxide, chromium oxide, tungsten oxide, nickel oxide or sulfide or any of a number of oxides or sulfides of metals of groups IV, V, VI, VII and VIII of the periodic system. These catalytic materials are usually dispersed or supported on a base or spacing agent. A commonly used spacing agent for this type of catalyst is alumina either precipitated or of the gel type. Catalysts which are more heat stable and which give favorable yields of aromatic gasoline components, particularly with certain feeds and at relatively low pressures, have been prepared upon zinc aluminate spinel supports.
It has been proposed in application Serial No. 188,236 filed October 3, 1950, now U. S. Patent 2,689,823, to effect the hydroforming of naphtha fractions in a fluidized solids reactor system in which naphtha vapors are passed continuously through a dense, fluidized bed of hydroforming catalyst particles in a reaction zone, spent catalyst particles being withdrawn from the dense bed in the reaction zone and passed to a separate regeneration zone where inactivating carbonaceous deposits are removed by combustion whereupon the regenerated catalyst particles are returned to the main reactor vessel. Fluid hydroforming as thus conducted has several fundamental advantages over fixed bed hydroforming such as (1) the operations are continuous, (2) the vessels and equipment can be designed for single rather than dual functions, (3) the reactor temperature is substantially constant throughout the bed and (4) the regeneration or reconditioning of the catalyst may be readily controlled.
A major part of the investment in a fluid hydroformer plant is for equipment required to supply the necessary heat of reaction to the reactor. In conventional plants, this heat of reaction is supplied as sensible heat from the inflowing feed, from the hydrogen-rich or recycle gas stream and from the recirculated catalyst stream. Difliculties are encountered when maximizing the heat input through these expedients since thermal degradation of the recycle gas to remove small amounts of C4 and C5 hydrocarbons must be resorted to or thermal degradation of these components will occur in the recycle gas preheating step and degradation or deactivation of the catalyst occurs if the temperature of the catalyst is allowed to exceed about 1200 F. in either the regenerator or in the reactivation of hydrogen treatment of the regenerated catalyst prior to recycling the same to the mainreactor bed.
It is the object of this invention to provide the art with an improved fluidized solids hydroforming process.
It is also the object of this invention to provide an improved fluidized solids hydroforming process which effects the introduction or supply of heat directly or indirectly to the reactor solids without sacrificing yield of liquid products.
These and other objects of this invention will appear more clearly from the detailed specification and claims which follow.
It has now been found that heat may be supplied directly or indirectly to the reactor solids in a fluid hydroforming system by providing a two-stage hydroforming reactor, transporting catalyst from a first reaction stage to a second stage by the use of the reactor product gases and recycling the catalyst from the second stage to the first stage through a catalyst heater in which the catalyst particles are brought to such temperature as to maintain the first reaction stage at the desired temperature. It is possible in this arrangement, to keep the catalyst to oil ratio between the reactor and the regenerator to a minimum thereby minimizing product losses associated with catalyst to oil ratios and to set the regenerator temperature well below the usual level of 1200 F. thereby minimizing loss in catalyst activity or life associated with high regenerator temperature or high temperature pretreat. It is also possible in this way to set the recycle gas rate at a value which will give maximum yield and since the preheat furnaces on both the oil feed and the recycle gas stream are eliminated, the investment and yield losses associated with these furnaces are eliminated. It further is possible in this arrangement to avoid the use of excess air and since the solids circulating through the heater are not oxidized, no yield loss results from their circulation.
Reference is made to the accompanying drawing in which:
Fig. 1 is a schematic flow plan of one embodiment of this invention;
Fig. 2 is a schematic flow plan of another embodiment in which the catalyst solids are heated by direct heat exchange with heated shot, and
Fig. 3 is a schematic flow plan of a further embodiment in which heat is supplied to the catalyst solids through heat transfer surfaces.
Referring to Fig. 1, 10 is a reactor Vessel which may which is in finely divided form and maintained as a dense, fluidized turbulent bed 11 by the passage therethrough of hydrogen-rich or recycle gas introduced through line 12 and hydrocarbon feed introduced through inlet line 13. A perforated plate or distribution grid 14 is preferably provided near the bottom of the reactor vessel in order to insure uniform distribution of the incoming constituents over the entire cross-section of the vessel. A distributor ring 15 is preferably provided on feed inlet line 13 in order to insure uniform distribution of the hydrocarbon feed.
The reactor vessel 10 is operated substantially full, i. c. with the dense bed 11 extending substantially to the top of the reactor so that catalyst is carried out by product gases through outlet line 16 into the second reactor vessel 20. The outlet line 16 terminates in an inlet cone 17 arranged in the lower part of the second reactor ves sel 20. The inlet cone is provided with perforations 18 to insure even distribution of the reaction mixture into vessel 20. catalyst particles in the second reactor vessel 20 in the form of a dense, fluidized, liquid-simulating bed 21. The bed 21 has a definite level L and is superposed by a dilute or disperse phase 22 comprising gaseous or vaporous reaction products containing a small amount of catalyst entrained therein. The reaction products are taken overhead from reactor vessel 29, preferably after passage through cyclone separators 23 which serve to knock out entrained catalyst particles which are returned to the dense catalyst bed 21 through the diplegs 24. The reaction products pass overhead through outlet line 25 to suitable fractionating, stabilizing and/or storage equipment.
Means are provided for the Withdrawal of a stream of catalyst directly from the dense bed 11 in order to regenerate the catalyst. This may be in the form of a cell or conduit 28 arranged entirely Within the reactor with its upper end 29 sealed and connected to the upper section of vessel 20 through a pipe. The purpose of this connection is to provide a means for discharging the stripper gases in the zone above the dense bed. One or more restricted inlet ports 30 control the passage of catalyst particles from bed 11 into conduit 28. The catalyst par ticles pass downwardly through the withdrawal conduit 23 in countercurrent to an inert stripping gas such as steam or the like introduced at 31. The catalyst particles are discharged from the withdrawal conduit 28 through a slide valve 32 in standpipe 33 into transfer line 34 where they are picked up by a stream of air or regeneration gas supplied at 35 and conveyed into regenerator vessel 36.
It may be desirable in some cases to utilize only part of the air or regeneration gas, generally not more than about 15 to of the total amount required for regeneration as carrier gas for conveying spent catalyst into the regenerator through transfer line 34 and to introduce the remainder of the regeneration gas directly into the regenerator. In order to insure uniform distribution of the incoming regeneration gas and catalyst particles over the entire cross section of the regeneration vessel it may be desirable to provide a perforated plate or distribution grid at the lower end of the regenerator vessel. The velocity of the regeneration gases through the vessel 36 is so controlled as to form a dense fluidized bed 37 having a definite level L superposed by a dilute or disperse phase 38. Regeneration gases are taken overhead through a cyclone separator 39 for removal of entrained catalyst which is returned to the dense bed 37 via the dipleg attached to the cyclone separator and passed through outlet line 40 to the recycle reactor catalyst heater as will be described in detail below.
Means are provided in the second reactor vessel 20 for the withdrawal of a stream of reactor catalyst for recycle to the first reactor vessel 10. This may be in the form of a cell or conduit 41 arranged entirely within the reactor vessel 20 and with its upper end above the dense bed level as in the case of withdrawal conduit 28 in reactor 10 The reaction product vapors maintain the or it may have its upper end 42 below the dense bed level L so that the catalyst overflows from the bed 21 into the withdrawal conduit 41 as shown.
The withdrawn catalyst passes downward through conduit 41 countercurrent to a stream of inert stripping gas such as steam, methane or the like introduced at 43 and through a slide valve or the like 44 for controlling the withdrawal of catalyst from reactor 20. The slide valve 44- may, if desired, be connected to a level controller 45 in order to maintain the dense bed level or reactor 20. The catalyst is discharged from the base of conduit 41 into catalyst heater 50.
The reactor catalyst particles are heated in heater by the flue gases from the regenerator 36 and, if necessary, by the combustion of additional fuel. Accordingly the flue gases withdrawn from regenerator 36 are passed through line 40 into transfer line 46 where they are mixed with additional air supplied through line 47. The mixture of regenerator flue gases and additional air is passed through line 46 to combustion chamber 48 where a fuel such as torch oil, polymer slurry formed in the process or excess process gas is supplied through line 49 in stoichiometric amounts. The quantity of air and fuel supplied in the burner depends upon the amount of heat that it is desired to generate. The hot combustion gases pass directly from combustion chamber 48 into heater 50 where they contact the catalyst particles supplied through conduit 41. The superficial velocity of the combustion or flue gases through heater 50 is so controlled as to form a dense, fluidized bed 51 having a definite level L". The flue gases pass overhead from heater 50, are separated from entrained solids in the conventional manner and are discharged through line 52 to a waste gas stack or the like. The heated catalyst particles overflow from dense bed 51 into standpipe 53 and through a slide valve 54 or other flow control means back into the dense bed 11 in reactor 10. Entrained flue gases are stripped from the solids in standpipe 53 by means of a suitable stripping gas such as nitrogen or by the use of tail gas produced in the process which is introduced at 55.
Regenerated catalyst, i. e. catalyst that has been sub stantially freed of carbonaceous deposits is withdrawn from regenerator 36 through outlet line 56 and discharged into treater 57. Hydrogen-rich treating gas is supplied to treater 57 through inlet line 58 and serves to partially reduce higher catalytic metal oxides formed during regeneration to more catalytically active compounds. Molyb denum oxide, for example, which is present in the reactor 10 as M0205 is oxidized in regenerator 36 to M003 and then reduces in treater 57 to M0205. The treating gases are discharged from treater 57 through line 59 while the treated or partially reduced catalyst particles are discharged through line 60 and slide valve 61 into the dense bed 11 in reactor 10.
Figure 2 shows essentially the same fiow plan and the several parts that are the same have been designated by the same reference numeral. This embodiment differs from that of Figure l merely by the fact that the contact solids or catalyst particles are heated in heater 50 by direct heat exchange with hot shot. In this embodiment, the flue gases withdrawn from regenerator 36 are passed through lines 40 and 46 to a distributor cone or the like 65 in the lower part of heater 50. The flue gases pass through the fluid bed 51 and are taken overhead as in the case of Fig. 1. Additional heat is supplied to the heater 50 by the circulation of shot therethrough. Catalyst is elutriated from the shot leaving vessel 50 through stripper S by means of a gas such as air, or process tail gas. By adjusting the gas velocity in S, the catalyst will be blow out the top of S and catalyst free shot discharged from the bottom of S. The shot is discharged through valve controlled line 66 into transfer line burner 67, where it is contacted with a burning mixture of fuel and air which serves to heat the shot to the desired elevated temperature. The heated shot is separated from the combustion products or flu gas in separator E, and then returned to the heater vessel 50 through valve controlled line 68. The inlet of catalyst discharge line 53 is set near the interface L" in vessel 50, so that the catalyst in vessel 50 overflows into line 53. Because of the overflow feature the catalyst entering line 53 will be free of the heavier shot.
Figure 3 also shows essentially the same flow plan as is shown in Figure 1 and therefore the several parts that are the same have been designated by the same reference numeral. This embodiment differs from that of Figure 1 merely by the fact that heat is supplied to the catalyst solids in heater 50 by indirect heat transferor through a heat exchanger surface. In this case, heat in addition to the heat supplied by the regenerator flue gas is supplied to the reactor catalyst in heater 50 through the heat exchanger 70. As heat exchange medium there is preferably used a stream of direct fired solids. In this case finely divided insert solid materials such as sand are charged to burner 71. Air and any desired fuel, solid, liquid or gaseous, are supplied to the burner 71 through line 72. Heat generated by the combustion of the fuel in burner 71 is absorbed by the inert solids in the burner vessel 71, the waste flue gases passing overhead through outlet line '73. The heated inert solids pass by gravity down through conduit 74 serving as a standpipe and through slide valve 75 or other similar flow control device into transfer line 76, where they are picked up by a stream of air or the like and conveyed into heat exchanger 70 within the heater 50. Control of the amount of heat transferred to the reactor catalyst solids in bed '51 in heater 50 may be achieved by control of the temperature of the solids in burner 71, by control of the rate of circulation of the hot inert solids through heat exchanger -70 and by control of the density and the residence time of the solids in the heat exchanger 70. The cooled heat transfer solids pass from heat exchanger 70 into transfer line 77 and thence back into the burner 71. A particular advantage of the system shown in Figure 3 is that the air for firing the burner 71 must be compressed to pressures only slightly above atmospheric pressure since the inert heat transfer solids are circulated through the system at substantially atmospheric pressure. In Figures 1 and 2, on the other hand, the air required for the generation of heat must be compressed to at least system pressure of about 50-1000 lbs. per sq. inch and ordinarily about 200 lbs. per sq. inch.
The feed or charging stock to the hydroforming reactor may be a virgin naphtha, a cracked naphtha, a Fischer- Tropsch naphtha or the like. The feed stock may be preheated as by indirect heat exchange with the hot reaction products passing through line and if desired may be given a further preheat before charging the same to the reactor. The recycle gas, which contains from about 50 to 80 vol. per cent hydrogen is preheated similarly to the naphtha feed stock by indirect heat exchange with reaction products and, if desired, it may be further preheated in a furnace or the like. It is noted that in view of the amounts of heat that can be introduced into the reactor zones by means of the catalyst particles in accordance with the present invention, preheating of the fresh feed and recycle gas may be minimized or eliminated if desired. The recycle gas is circulated through the reactor system at a rate of from about 1000 to 8000 cu. ft. per barrel of feed.
The reactor system is charged with a mass of finely divided hydroforming catalyst particles. Suitable catalysts include group VI metal oxides such as molybdenum oxide, chromium oxide or tungsten oxide or mixtures thereof alone or preferably supported upon a base or carrier such as activated alumina, zinc aluminate spinel or the like. Prefered catalysts contain about 5 to 15 weight percent molybdenum oxide or about 10 to 40 weight percent chromium oxide upon a suitable carrier. If desired, minor amounts of stabilizers or promoters such as silica, calcium oxide, ceria or potassia can be included in the catalyst. Catalysts comprising platinum and/or palladium upon a support such as activated alumina or upon alumina gel may also be used in accordance with the present invention. These alumina supports may contain small amounts of silica, generally up to about 6.0 weight per cent and the catalyst may also contain added halogen, for example fluoride ion in amounts of 0.1 to 3.0 weight per cent as by treating the support, before or after depositing platinum or palladium thereon, with hydrogen fluoride. The platinum or palladium may be incorporated on the base by impregnating the same with an aqueous solution of a water-soluble compound of the catalytic metal, drying and calcining the same or treating the impregnated base with a precipitant for the catalytic metal compound, such as hydrogen sulfide, drying and calcining the composite. Before placing the catalyst in use in a reforming operation, it is advisable to treat the catalyst with hydrogen or hydrogen-rich gas at elevated temperatures. This may be effected by treating the catalyst with hydrogen or hydrogen-rich gas in the reforming reactor unit while gradually raising its temperature to about 800-1000 F. Effective catalysts of this type contain from about 0.2 to 1.0 weight per cent, preferably about 0.5 weight per cent of platinum or from about 0.5 to 2.0 weight per cent, preferably about 1.0- 1.5 weight per cent of palladium. For proper fluidization, the catalyst particles should, for the most part, be between about 200 and 400 mesh in size or about 0-200 microns in diameter with a major proportion between 20 and microns.
The hydroforming reactor vessel 10 should be operated at temperatures between about 850 and 925 F., preferably about 900 F. and at pressures of between 50 and about 500 lbs. per sq. inch, preferably about 200 lbs. per sq. inch. The second reactor vessel 20 is operated at a slightly lower temperature, generally about 5 to 10 F. below the temperature of the first reactor vessel 10. Reactor temperatures above about 900 F. generally result in increased carbon formation and lower selectivity to gasoline fractions while at temperatures below 900 F., operating severity is low and would therefore require an excessively large reaction vessel. Lowering reactor pressure below 200 lbs. per sq. inch ordinarily results in increased carbon formation which becomes excessive in most cases at pressures below about 75 lbs. per sq. inch. Above 200 lbs., however, catalyst selectivity to light products (C4s and lighter) increases rapidly. The regenerator vessel is normally operated at essentially the same pressure as the reactor vessels and at temperatures of about 1000-1200 F. The residence time of the catalyst in the reactor vessels is of the order of from 2 to 4 hours and in the regenerator of from about 3 to 15 minutes.
The weight ratio of catalyst to oil introduced into the reactor should be about 0.5 to 1.5. It is preferred to operate at catalyst-to-oil ratios of about 1 since ratios above about 1 to 1.5 result in excessive carbon formation. Somewhat higher weight ratios can be used at higher pressures. Space velocity or the weight in pounds of feed charged per hour per pound of catalyst in the reactor depends upon the age or activity level of the catalyst, the character of the feed stock and the desired octane number of the product. Space velocity for a molybdenum oxide on alumina gel catalyst may vary, for example from about 1.5 wt./hr./wt. to about 0.15 wt./hr./wt.
The treatment of the freshly regenerated catalyst with hydrogen-rich gas in the treater 57 is an exothermic reaction and care should be taken that the temperature of the catalyst undergoing pretreatment does not exceed about 1200 F. In general, the catalyst particles should not be in contact with hydrogen containing gas at temperatures above about 1150 F. for more than a few seconds (less than 15 and preferably less 5 seconds). At lower temperatures, for example, 1050 F. and below,
time of contact of the freshly regenerated catalyst and the hydrogen-rich gas can be for as long as 15 minutes without causing degradation of the catalyst.
The foregoing description contains a limited number of embodiments of the present invention. It will be understood, however, that numerous modifications may be made by those skilled in the art without departing from the scope of this invention.
What is claimed is:
1. In a process for reforming hydrocarbons in contact with finely divided reforming catalyst particles, the process which comprises continuously supplying catalyst particles, hydrocarbon feed stock and hydrogen-rich gas to a primary reaction zone maintained at reforming temperatures and pressures, taking a stream of catalyst particles in reactant vapors overhead from the first reaction zone, passing said stream into a second reaction zone, controlling the velocity of the reactant vapors through the second reaction zone to form a dense, fluidized bed of catalyst particles suspended in reactant vapors in said second reaction zone, withdrawing a stream of reaction products overhead from said second reaction zone, Withdrawing a stream of catalyst particles downwardly directly from the dense bed in said second reaction zone, heating this stream of catalyst particles in a non-oxidizing atmosphere to a temperature substantially above the temperature in the reaction zones, recycling the stream of heated catalyst particles to the first reaction zone, withdrawing a second stream of catalyst particles from the first reaction zone, passing this second stream of catalyst particles to a regeneration zone, removing carbonaceous deposits from the catalyst particles in the regeneration zone, and recycling the regenerated catalyst to the first reaction zone.
2. In a process for reforming hydrocarbons in contact with finely divided reforming catalyst particles comprising platinum upon a support, the process which comprises continuously supplying catalyst particles, hydrocarbon feed stock and hydrogen-rich gas to a primary reaction zone maintained at reforming temperatures and pressures, taking a stream of catalyst particles in reactant vapors overhead from the first reaction zone, passing said stream into a second reaction zone, controlling the velocity of the reactant vapors through the second reaction zone to form a dense, fluidized bed of catalyst particles suspended in reactant vapors in said second reaction zone, withdrawing a stream of reaction products overhead from said second reaction zone, withdrawing a stream of catalyst particles downwardly directly from the dense bed in said second reaction zone, heating this stream of catalyst particles in a non-oxidizing atmosphere to a temperature substantially above the temperature in the reaction zones, recycling the stream of heated catalyst particles to the first reaction zone, withdrawing a second stream of catalyst particles from the first reaction zone, passing this second stream of catalyst particles to a regeneration zone, removing carbonaceous deposits from the catalyst particles in the regeneration zone by treating the same with hydrogen-rich gas and recycling the hydrogen treated regenerated *catalyst to the first reaction zone.
3. In a process for reforming hydrocarbons in contact with finely divided catalyst particles comprising a group VI metal oxide upon a support, the process which comprises continuously supplying catalyst particles, hydrocarbon feed stock and hydrogen-rich gas to a primary reaction zone maintained at hydroforming temperatures and pressures, taking a stream of catalyst particles in reactant vapors overhead from the first reaction zone, passing said stream into a second reaction zone, controlling the velocity of the reactant vapors through the second reaction zone to form a dense, fluidized bed of catalyst particles suspended in reactant vapors in said second reaction zone, withdrawing a stream of reaction products overhead 'from said second reaction zone, withdrawing a stream of catalyst particles downwardly directly from the dense bed in said second reaction zone, heating this stream of catalyst particles in a non-oxidizing atmosphere to a temperature substantially above the temperature in the reaction zones, recycling the stream of heated catalyst particles to the first reaction zone, withdrawing a second stream of catalyst particles from the first reaction zone, passing this second stream of catalyst particles to a regeneration zone, burning carbonace'ou's deposits from the catalyst particles in the regeneration zone, treating the regenerated catalyst particles with hydrogen-containing gas and recycling the hydrogentreated regenerated catalyst to the first reaction zone.
4. The process as defined in claim 3 in which the hot combustion gases formed in the regeneration zone are utilized for heating the stream of catalyst particles withdrawn from the second reaction zone.
5. The process as defined in claim 3 in which the hot combustion gases formed in 'the regeneration zone are depleted of oxygen by burning an auxiliary fuel therein and the resultant hot'gases are utilized for heating the stream of catalyst particles withdrawn from the second reaction zone.
6. The process as defined in claim 3 in which the stream of catalyst particles withdrawn from the second reaction zone is heated by 'intermixing hot, inert, heat transfer solids therewith whereupon the heated catalyst particles are returned to the first reaction zone and the inert heat transfer solids are transferred to a heating zone preparatory to recycling the same.
References'Cited in the file of this patent UNITED STATES PATENTS 2,399,050 Martin Apr. 23, 1946 2,410,891 Meinert et al Nov. 12, 1946 2,472,844 Munday et al June 14, 1949 2,515,156 Jahnig et a1. July 11, 1950 2,540,373 McAfee Feb. 6, 1951 2,602,771 Munday July 8, 1952

Claims (1)

1. IN A PROCESS FOR REFORMING HYDROCARBONS IN CONTACT WITH FINELY DIVIDED REFORMING CATALYST PARTICLES, THE PROCESS WHICH COMPRISES CONTINUOUSLY SUPPLYING CATALYST PARTICLES, HYDROCARBON FEED STOCK AND HYDROGEN-RICH GAS TO A PRIMARY REACTION ZONE MAINTAINED AT REFORMING TEMPERATURES AND PRESSURES, TAKING A STREAM OF CATALYST PARTICLES IN REACTANT VAPORS OVERHEAD FROM THE FIRST REACTION ZONE, PASSING SAID STREAM INTO A SECOND REACTION ZONE, CONTROLLING THE VELOCITY OF THE REACTANT VAPORS THROUGH THE SECOND REACTION ZONE TO FORM A DENSE, FLUIDIZED BED OF CATALYST PARTICLES SUSPENDED IN REACTANT VAPORS IN SAID SECOND REACTION ZONE, WITHDRAWING A STREAM OF REACTION PRODUCTS OVERHEAD FROM SAID SECOND REACTION ZONE, WITHDRAWING A STREAM OF CATALYST PARTICLES DOWNWARDLY DIRECTLY FROM THE DENSE BED IN SAID SECOND REACTION ZONE, HEATING THIS STREAM OF CATALYST PARTICLES IN A NON-OXIDIZING ATMOSPHERE TO A TEMPERATURE SUBSTANTIALLY ABOVE THE TEMPERATURE IN THE REACTION ZONES, RECYCLING THE STREAM OF HEATED CATALYST PARTICLES TO THE FIRST REACTION ZONE, WITHDRAWING A SECOND STREAM OF CATALYST PARTICLES FROM THE FIRST REACTION ZONE, PASSING THIS SECOND STREAM OF CATALYST PARTICLES TO A REGENERATION ZONE, REMOVING CARBONACEOUS DEPOSITS FROM THE CATALYST PARTICLES IN THE REGENERATION ZONE, AND RECYCLING THE REGENERATED CATALYST TO THE FIRST REACTION ZONE.
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US2815268A (en) * 1953-08-05 1957-12-03 Exxon Research Engineering Co Apparatus for treating and separating finely divided solids
US2846364A (en) * 1954-12-01 1958-08-05 Exxon Research Engineering Co Fluid hydroforming process using a platinum catalyst and inert heat transfer solids
US2866747A (en) * 1953-03-04 1958-12-30 Exxon Research Engineering Co Aromatization and naphtha reforming process wherein the catalyst is dehydrated with a gas containing c1-c4 hydrocarbons
US2883335A (en) * 1953-07-01 1959-04-21 Kellogg M W Co Hydrocarbon conversion system
US2884368A (en) * 1956-02-17 1959-04-28 United Eng & Constructors Inc Process for the pyrolysis and gasification of hydrocarbonaceous materials
US2890168A (en) * 1953-10-12 1959-06-09 Kellogg M W Co Hydrocarbon conversion system
US2905622A (en) * 1954-04-29 1959-09-22 Phillips Petroleum Co Production of fuel gas and liquid hydrocarbon fuels
US2906697A (en) * 1954-10-11 1959-09-29 Gulf Research Development Co Hydroreforming and reactivation of catalyst with an oxygen containing gas
US2909480A (en) * 1956-01-18 1959-10-20 Gulf Research Development Co Hydroreforming process
US2934494A (en) * 1957-08-15 1960-04-26 Exxon Research Engineering Co Recovery of finely divided solids
US2957922A (en) * 1956-09-28 1960-10-25 Diamond Alkali Co Chlorination of ethylene dichloride
US3007862A (en) * 1958-09-17 1961-11-07 Kellogg M W Co Hydrocarbon conversion system
US3024186A (en) * 1958-03-18 1962-03-06 British Petroleum Co Hydroforming
US3033906A (en) * 1958-08-06 1962-05-08 Gulf Research Development Co Process for converting normal hexane to benzene
US3111480A (en) * 1958-03-31 1963-11-19 Socony Mobil Oil Co Inc Sequential high pressure-low pressure reforming
US3169916A (en) * 1953-05-19 1965-02-16 Standard Oil Co Multistage hydrocarbon reforming with fluidized platinum catalyst
US4725408A (en) * 1984-04-24 1988-02-16 Texaco, Inc. Fluid catalytic cracking apparatus

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US2765260A (en) * 1952-08-01 1956-10-02 Exxon Research Engineering Co Hydroforming process and apparatus
US2866747A (en) * 1953-03-04 1958-12-30 Exxon Research Engineering Co Aromatization and naphtha reforming process wherein the catalyst is dehydrated with a gas containing c1-c4 hydrocarbons
US3169916A (en) * 1953-05-19 1965-02-16 Standard Oil Co Multistage hydrocarbon reforming with fluidized platinum catalyst
US2883335A (en) * 1953-07-01 1959-04-21 Kellogg M W Co Hydrocarbon conversion system
US2815268A (en) * 1953-08-05 1957-12-03 Exxon Research Engineering Co Apparatus for treating and separating finely divided solids
US2890168A (en) * 1953-10-12 1959-06-09 Kellogg M W Co Hydrocarbon conversion system
US2905622A (en) * 1954-04-29 1959-09-22 Phillips Petroleum Co Production of fuel gas and liquid hydrocarbon fuels
US2906697A (en) * 1954-10-11 1959-09-29 Gulf Research Development Co Hydroreforming and reactivation of catalyst with an oxygen containing gas
US2846364A (en) * 1954-12-01 1958-08-05 Exxon Research Engineering Co Fluid hydroforming process using a platinum catalyst and inert heat transfer solids
US2909480A (en) * 1956-01-18 1959-10-20 Gulf Research Development Co Hydroreforming process
US2884368A (en) * 1956-02-17 1959-04-28 United Eng & Constructors Inc Process for the pyrolysis and gasification of hydrocarbonaceous materials
US2957922A (en) * 1956-09-28 1960-10-25 Diamond Alkali Co Chlorination of ethylene dichloride
US2934494A (en) * 1957-08-15 1960-04-26 Exxon Research Engineering Co Recovery of finely divided solids
US3024186A (en) * 1958-03-18 1962-03-06 British Petroleum Co Hydroforming
US3111480A (en) * 1958-03-31 1963-11-19 Socony Mobil Oil Co Inc Sequential high pressure-low pressure reforming
US3033906A (en) * 1958-08-06 1962-05-08 Gulf Research Development Co Process for converting normal hexane to benzene
US3007862A (en) * 1958-09-17 1961-11-07 Kellogg M W Co Hydrocarbon conversion system
US4725408A (en) * 1984-04-24 1988-02-16 Texaco, Inc. Fluid catalytic cracking apparatus

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