US3870481A - Method for production of synthetic natural gas from crude oil - Google Patents

Method for production of synthetic natural gas from crude oil Download PDF

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US3870481A
US3870481A US297012A US29701272A US3870481A US 3870481 A US3870481 A US 3870481A US 297012 A US297012 A US 297012A US 29701272 A US29701272 A US 29701272A US 3870481 A US3870481 A US 3870481A
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hydrogen
crude oil
methane
effluent
ethane
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William P Hegarty
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Priority to US297012A priority Critical patent/US3870481A/en
Priority to CA181,269A priority patent/CA1001416A/en
Priority to ZA737505*A priority patent/ZA737505B/xx
Priority to AU60657/73A priority patent/AU463986B2/en
Priority to IN2183/CAL/1973A priority patent/IN140810B/en
Priority to GB4512973A priority patent/GB1453081A/en
Priority to GB1705876A priority patent/GB1453082A/en
Priority to JP48113167A priority patent/JPS4970901A/ja
Priority to DE19732350666 priority patent/DE2350666A1/de
Priority to ES419508A priority patent/ES419508A1/es
Priority to FR7336265A priority patent/FR2208967B1/fr
Priority to IT53024/73A priority patent/IT996290B/it
Priority to BE2053138A priority patent/BE805927A/xx
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    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10L—FUELS NOT OTHERWISE PROVIDED FOR; NATURAL GAS; SYNTHETIC NATURAL GAS OBTAINED BY PROCESSES NOT COVERED BY SUBCLASSES C10G, C10K; LIQUEFIED PETROLEUM GAS; ADDING MATERIALS TO FUELS OR FIRES TO REDUCE SMOKE OR UNDESIRABLE DEPOSITS OR TO FACILITATE SOOT REMOVAL; FIRELIGHTERS
    • C10L3/00—Gaseous fuels; Natural gas; Synthetic natural gas obtained by processes not covered by subclass C10G, C10K; Liquefied petroleum gas
    • 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
    • C10G9/00—Thermal non-catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • Y02P30/00—Technologies relating to oil refining and petrochemical industry
    • Y02P30/40—Ethylene production

Definitions

  • ABSTRACT A process for the production of a pipeline gas of high BTU content from crude oil by hydrogasification of the crude oil.
  • the crude oil is first vaporizedin the presence of hydrogen and then gasified to form an effluent gas containing essentially methane, ethane, aromatic hydrocarbons, hyrogen and hydrogen sulfide. After separation of the aromatics and hydrogen sulfide from the effluent, the effluent is subjected to cryogenic separation of the hydrogen and a final catalytic conversion of the ethane to methane.
  • This invention pertains to the production of synthetic natural gas (methane) by the gasification of crude oil.
  • Crude oil is generally separated into high and low boiling point fractions and the low boiling point fractions are then treated at high temperature to produce an effluent gas containing methane, ethane, acid gases such as hydrogen sulfide, excess hydrogen, and residual aromatic constituents.
  • the effluent from the gasification step is then subjected to further processing sequences where the acid gas and aromatic fractions are removed and the hydrogen is separated, and finally the ethane is reacted to form additional methane and the methane or synthetic natural gas is fed to a pipeline for use in, among other things, residential communities and industrial concerns.
  • One commercial process technique for producing synthetic natural gas by hydrocarbon gasification of naphtha distillate feed stocks is the stream reforming process operated either under severe or mild reforming conditions.
  • the steam reforming process is carried out at elevated temperatures at atmospheric or higher pressures in the presence of a catalyst, generally a supported nickel catalyst.
  • Hydrocarbons react with the steam to give a gas comprising hydrogen, carbon oxide, and methane with the composition of the gas depending upon the conditions of the reaction.
  • Steam reforming under severe conditions is disclosed in U.S. Pat. Nos. 3,103,423 and 3,063,936.
  • British Pat. No. 981,726 discloses a process for steam reforming under mild conditions wherein a methane rich gas is produced. This is illustrative of what is known in the trade as the British Gas Council Catalytic Rich Gas Process (CRG Process).
  • CCG Process British Gas Council Catalytic Rich Gas Process
  • a second method of producing a synthetic pipeline gas, and one that is especially rich in methane, is by the various thermal hydrocracking techniques.
  • Thermal hydrocracking can be broken down into various categories. The first of these is the operation of the thermal hydrocracking unit without a coke bed or a catalyst; the second is operation with a fluidized coke bed in order to control carbon deposition by periodic withdrawal of solids; and the last is the operation with a hydroforming catalyst to influence the reaction.
  • the reaction remains largely thermal hydrocracking rather than catalytic hydrocracking at low pressures and high temperatures involved with steam being used to control coke deposition (lay down) on the catalyst.
  • thermal hydrocracking processes are exemplified in U.S. Pat. Nos. 2,759,806, 2,882,138, 2,926,077, 3,124,436, 3,202,603, 3,484,219, and 3,591,356.
  • the foregoing Patents are illustrative of thermal hydrocracking processes that variously employ a fluidized bed, recirculation of the gas, or a multiple stage gasification system.
  • thermal hydrocracking processes or hydrocracking processes accomplished without the influence of a catalyst will produce large quantities of methane in the production of synthetic pipeline gases.
  • Hydrocarbon feed stocks, boiling over a wide range can be used as the starting product for conventional thermal hydrocracking processes.
  • the conventional processes can operate at temperatures of between 750 and 1650F.
  • reactions include thermal cracking and hydrogenation, and production of a large volume of light gases including methane is characteristic in large part because of the required high temperature to produce cracking in the absence of a catalyst. In such processes, steam may be added to the reaction zone in order to control the deposition of coke on the vessel interior.
  • GSH Process Gas Recycle Hydrogenator Process
  • This process has been used successfully in producing methane from light petroleum distillates, which can be subsequently processed to a synthetic natural gas for pipeline use.
  • the process is basically a thermal hydrocracking one and is conducted at temperatures of between l290 and 1470F, and at pressures on the order of from to 1350 psig in the presence of a hydrogen gas.
  • the hydrogen gas reacts with the hydrocarbon feed in an exothermic reaction to produce gaseous hydrocarbons; mainly, methane and ethane.
  • the hydrogen gas can be either pure hydrogen or a gas consisting predominately of hydrogen, such as the product gas from a partial oxidation unit or a steam reforming unit.
  • a third category of gasification processes which under certain conditions may be used in the process of the present invention include the partial oxidation process.
  • a gaseous or liquid hydrocarbon feed is partially oxidized using oxygen, or air with light feeds, or steam with heavy feeds.
  • Reactors are operated at pressures of to 300 psig at temperatures of between 1800 and 3500F. Such reactors produce predominantly hydrogen and carbon monoxide and are well known in the art.
  • Two commercial processes that are offered for sale have been developed by Shell and Texaco. The Shell process is shown in U.S. Pat. No. 2,971,829.
  • a fourth process for the manufacture of fuel gas is disclosed in U.S. Pat. No. 3,531,267 and combines a catalytic hydrocracking and gasification process.
  • an improved process for producing a pipeline gas of high heating value from a crude oil comprises basically the steps of vaporizing a substantial portion of the crude oil at a temperature of between 600 and 1000F, thereafter introducing the vaporized crude oil and a hydrogenation gas into a gasification vessel maintained at a temperature in excess of 1000F wherein the feed stream is gasified producing an effluent consisting essentially of hydrogen, hydrogen sulfide, methane, ethane, and residual aromatic hydrocarbons, thereafter cooling the effluent gas to room temperature and recovering the waste heat, drying the effluent and removing the hydrogen sulfide and residual aromatics from the effluent, cryogenically separating the methane and ethane from the hydrogen, and thereafter reacting the ethane with steam to produce additional methane and carbon dioxide, and removing the carbon dioxide.
  • the two methane streams are combined and discharged into a process pipeline or storage vessel.
  • the methane from the gasifier does not have to be separated from the ethane during ethane conversion.
  • the process includes additional steps wherein the various side streams can be used to produce sources of reactant or heat for operating the plant utilities in order to achieve an efficient overall process. While it would be possible to use various hydrogasification methods, the gas recycle hydrogenator developed by the British Gas Council is the preferred hydrogenation vessel for the method of the instant application.
  • the hydrogen separated from the methane and ethane can be recycled to the feed stream vaporization unit or the gasifier.
  • a key to the present invention is the vaporization step, which is also disclosed, comprising heating the oil and the hydrogenation gas, preferably hydrogen, and admixing the heated oil and gaseous hydrogen in a vaporizer and taking the vaporizer effluent and passing it in to the hydrogenation vessel, thereafter recovering the heat from the gasifier effluent and using a portion of the gasifier effluent after cooling to control the temperature of reaction of the gasification vessel thereby preventing unwanted coke formation in the gasification vessel.
  • the vaporization step comprising heating the oil and the hydrogenation gas, preferably hydrogen, and admixing the heated oil and gaseous hydrogen in a vaporizer and taking the vaporizer effluent and passing it in to the hydrogenation vessel, thereafter recovering the heat from the gasifier effluent and using a portion of the gasifier effluent after cooling to control the temperature of reaction of the gasification vessel thereby preventing unwanted coke formation in the gasification vessel.
  • FIG. 1 is a schematic drawing of the overall process for producing a synthetic natural gas from a crude oil.
  • FIG. 1a is a schematic drawing of an alternative method of converting naphtha and ethane to methane and carbon dioxide.
  • FIG. 2 is a schematic diagram of the improved method of feeding the hydrogasification vessel in order to achieve an effluent that is convertible to synthetic natural gas.
  • FIG. 3 is a schematic diagram of the test system used to verify the feeding and hydrogasification steps of the present invention.
  • FIG. 1 An overall process flow sheet for the method of the present invention.
  • numeral 10 indicates a feed stream preparation step in the overall process wherein a crude oil stream 12 is subjected to preliminary separation wherein a naphtha fraction having 360F end point is taken overhead through conduit 18 and used in a subsequent processing step as will hereinafter be more fully explained.
  • the process stream 15 after preparation may be subjected to a vaporization process as will be discussed in connection with FIG. 2, after which said stream comprises crude oil material having a +360F boiling point and hydrogenation gas 13 which is injected into a single-stage gasifier such as that disclosed in US. Pat. No. 3,363,024.
  • the vaporizer and/or gasifier is represented in'FlG. 1 by block 16.
  • the gasifier In the gasifier, about 83 percent of the oil is gasified at high temperature to a mixture of methane, ethane, excess hydrogen, hydrogen sulfide and residual aromatics. The 17 percent residue of the oil plus a portion of the aromatics separated from the gas after cooling are fed to a hydrogen plant 20.
  • hydrogen represented by arrow 13
  • Oxygen for the partial oxidation portion of the hydrogen plant 20 is supplied by a separate oxygen plant 24, which plants are commercially available.
  • the balance of the hydrogen plant 20 includes waste heat recovery systems, water gas shift, acid gas removal and methanation systems to produce high-purity hydrogen by the reaction of CO and steam and subsequent removal of hydrogen sulfide and C0
  • the CO from the hydrogen plant 20 is vented through conduit 26 and the hydrogen sulfide is passed bia a conduit 28 to a sulfur plant 30 wherein it is combined with the hydrogen sulfide from the purification unit 34 and removed from the system by the claus process as elemental sulfur as at 32.
  • the gasifier effluent stream 22 is cooled and passed through a purification section 34 wherein the gas is dried and the acid gases as well as benzene and other residual aromatics are removed.
  • Acid gases are taken to mean CO and hydrogen sulfide.
  • the benzene and other residual aromatics are removed through conduit 36 and conducted through conduits 38 and 40 to the conduit 42, which receives the residual oil fraction from the gasifier and the mixture introduced into the hydrogen plant for production of hydrogen by the partial oxidation process. A portion of the benzene and residual aromatics can be used as a source of fuel for operating the plant.
  • Hydrogen sulfide is removed through conduit 44 and sent to the sulfur plant 30.
  • the effluent-from purification step 34 in conduit 46 contains essentially hydrogen, methane and ethane and is sent to a cryogenic separation unit 48 wherein the process stream is cooled, the methane and ethane are liquefied, separated from the hydrogen, vaporized, rewarmed compressed and removed through conduit 50, and the hydrogen through conduit 52.
  • the hydrogen in conduit 52 is warmed to ambient and recycled to the gasifier or the feed stream preparation section 16 and respectively.
  • the combined methane-ethane liquid stream in conduit 50 is charged to a catalytic rich gas process unit 54 along with the naphtha that is taken from the feed stream preparation stage 10 via conduit 18.
  • the naphtha is desulfurized and the naphtha and ethane are reacted with steam in an autothermic catalytic reaction to produce methane plus carbon dioxide.
  • the carbon dioxide is separated and vented through conduit 56 and the methane product in conduit 58 is dried and compressed to 1000 psig and discharged into a product pipeline or suitable storage receptacle.
  • the methane produced in the gasifier is carried along with the process stream but does not enter into the overall reaction.
  • the methane from the gasifier can be separated before the catalytic rich gas unit 54 thereby reducing the overall size of unit 54.
  • FIG. 1a Another method of converting the naphtha ethane is shown in FIG. 1a wherein the naphtha from conduit 18 is reacted with steam 19 in a first catalytic rich gas reactor 54a thereby discharging steam, methane and carbon dioxide in conduit 160.
  • the methane and ethane from the cryogenic separation unit 48 are injected in conduit 160 through conduit 50 and the mixture is injected into a second catalytic rich gas reactor 5412 wherein the ethane and steam are reacted to form essentially methane and carbon dioxide.
  • the effluent in conduit 162 consists essentially of methane, carbon dioxide, carbon monoxide,
  • This effluent is sent to a methanation unit wherein additional methane is produced and the product gas is then sent to a carbon dioxide removal unit 166 through conduit 168. After carbon dioxide removal, the effluent gas consisting of about 94% methane with hydrogen and a trace of carbon monoxide is sent into a product pipeline through conduit 170.
  • a plant embodying such stages is completely self-sustaining regarding utilities.
  • by-product steam is generated in the hydrogen plant, the sulfur plant, and the gasification plant.
  • This steam is supplemented with an auxiliary steam generating plant to meet the requirements of the oxygen plant, the purification process, cryogenic separation process and the catalytic rich gas section.
  • Overall electrical power usage is neglibible since compressors and pumps are driven by condensing steam turbines.
  • Fuel for auxiliary steam generators is supplied by the residual aromatics from the process, which are burned as a fuel. It is calculated that overall thermal efficiency (product gas heating value as a percent of feed oil heating value) is about 84 percent.
  • the overall process as set forth in connection wit FIG. 1 comprises eight separate steps each of which is within the perview of current technology; however, such a combination has never before been shown or suggested.
  • the hydrogasification step 16 has not been successful for heavier crude oil (higher boiling feed stocks).
  • Hydrogasification is basically a highpressure, high-temperature, non-catalytic gas phase hydrocracking reaction. High molecular weight hydrocarbons are cracked and the fragments are saturated with hydrogen. Paraffins and naphthenes (cycloparaffins) are completely gasified to methane and ethane in the hydrogasification reaction alkyl side chains of the aromatics are gasified leaving residual benzene. Polycyclic aromatics are gasified leaving residual benzene.
  • GSH gas recycle hydrogenerator
  • US. Pat. No. 3,363,024 is the only hydrogen gasification process proven in cornmercail operation. This process features an adiabatic back-mixed reactor that dealkylates feed aromatics and completely gasifies paraffins and naphthenes (cycloparaffins) with a residence time of between 10 and 50 seconds at temperatures in the order of 1290F to 1390F under operating pressures of from to 1350 psig.
  • the product gas from such a gas recycle hydrogenator consists of methane, ethane, residual refractory aromatics and excess hydrogen.
  • the British Gas council has published its findings claiming that the only limitation on liquid feed stock is the ability to evaporate it as a gasifier feed.
  • Commercial applications have used naphthas with end points up to 365F.
  • the Gas Council has conductd pilot-plant tests with 635F end point gas oil with no gasifier carbon problems. However, in the latter case, plugging problems did develop and the Gas Council has accepted 650F as the maximum feeding end point for the process.
  • the Gas Council has identified three types of carbon depositions that can be found in the gas recycle hydrogenator. These are catalytic carbon, which is deposited on catalytically active metal surfaces with attendant pitting corrosion of the metal surface; wall carbon, which deposits on a surface of low activity already coated with carbon; and space carbon, which is produced as soot in the gas phase and exhausted with the gas product (effluent). Carbon formation is promoted by excessively high operating temperatures, carbon monoxide and carbon dioxide content and low gasifier hydrogen partial pressure. On the'other hand, carbon formation is suppressed by sulfur and steam, with a high internal recirculation rate to give a closer approach to complete back-mixing in uniform temperatures also tending to decrease carbon formation.
  • the hydrocarbon liquid/- hydrogen feed ratio must be maintained within certain limits. At high ratios, hydrogen partial pressure is low and the hydrogenation rate decreases with resulting cyclization of the reaction thereby causing a net production of aromatics. In an extreme case, condensation to polycyclic aromatics could result in tar and coke problems. Operation of the reactor includes maintaining the temperature by regulating the feed preheat temperatures. At low ratios of hydrocarbon liquid to hydrogen feed less reactant is charged, therefore, the heater reaction is less and the feed preheat temperatures increase. If such increases are excessive, the reaction could initiate in the preheater with attendant difficulties therein.
  • the Gas Council has reported its efforts were directed to getting the highest possible heating value gas direct from the gas recycle hydrogenator. Therefore, they sought to minimize the excess hydrogen thereby having a high ethane concentration in the effluent gas (about 0.5 mole per mole of methane). They have not reported any investigations using excess hydrogen.
  • One reported high-pressure run at I350 psia did hydrogenate ethane but the control was sluggish and operating temperature was unstable and carbon formation resulted in the vessels. If ethane hydrogenolysis is suddenly initiated with minimum excess hydrogemthere would be a sudden decrease in hydrogen partial pressure and the reacted temperature would rise. Such a combination would most surely produce carbon particularly when poor temperature control is present. Maintaining temperature control in the reactor with sufficient excess hydrogen to suppress carbon formation the additional heat release would decrease the required preheat temperature for the 'products being fed to the reactor.
  • the gas recycle hydrogenator is essentially a back-mixed reactor with the preheated feed (approximately -1000F) warmed to gasifier effluent temperature (about 1380F) almost instantaneously.
  • feed oil vapor concentrations are diluted to outlet concentrations immediately thereby minimizing self-condensation reactions; thus, formation of precursors at the intermediate temperature and high feed concentration levels should be negligible and tar and coke production minimized.
  • the gas recycle hydrogenator can accept feed stocks with higher boiling fractions when such feed stocks are gasified as shown in FIG. 2.
  • the basic gasification system is shown in FIG. 2 and consists of a vaporizer 60, gasifier 62, a liquid aromatics separator 64, a product gas compressor 66, together with hydrogen and oil heaters 68 and 70 respectively, and a gasifier product cooler 72.
  • a vaporizer 60 gasifier 62
  • a liquid aromatics separator 64 liquid aromatics separator 64
  • a product gas compressor 66 product gas compressor
  • hydrogen and oil heaters 68 and 70 respectively
  • gasifier product cooler 72 a gasifier product cooler
  • Such a system would operate at 600 psig or higher with the crude oil feed 74 and the hydrogen feed 76, each being preheated and charged to the vaporizer 60.
  • the oil 78 is further heatedby heater 80, which is submerged in the pool of liquid residual crude oil or heel as it is called in the trade.
  • the hydrogen after being preheated is sparged beneath the surface of the liquid residual crude oil 78 in vaporizer 60 so that the oil is vaporized into a mixture with the hydrogen and taken
  • the entering feed jet 86 induces a high internal recirculation rate in the gasifier along the path shown by the arrows.
  • the contents of the entering jet 86 are essentially completely mixed and have substantially uniform composition and temperature.
  • the temperature in the reactor is approximately 1400F and the average reactor residence time is between 10 and seconds.
  • the hot effluent gases are removed through conduit 88 and cooled in waste heat boiler 72 with heat recovery. After heat recovery, the effluent is conducted to aromatic separator 64 through a conduit wherein the condensed aromatic constituents are removed from the separator through conduit 92 and are used for supplementary feed to the partial oxidation unit and process fuel.
  • the resulting gas product from the separator 64 is compressed in compressor 66 and pushed on to further processing.
  • the gas at the exit end of the compressor consists essentially of hydrogen, methane, ethane, hydrogen sulfide, and uncondensed aromatics. This product can be further processed in accordance with the process described in FIG. 1.
  • a portion of the cooled gas from the compressor 66 is taken through conduit 94 and recycled through valve 96 into the reactor or hydrogasifier 62 to provide cooling of the reactor to maintain the reacted temperature at the desired temperature level of about 1400F.
  • FIG. 3 The foregoing process was verified in a test setup shown in FIG. 3 that was constructed and operated.
  • This system consisted of a hydrogen supply conduit 100, a hydrogen metering device (rotameter 102), electrically heated hydrogen humidifier 104, an oil feed system 106 comprising an oil tank 108, and oil metering pump 110, a vaporizer 112 that was electrically heated, a hydrogasification vessel 114, also electrically heated, a product cooler-condenser 116, a liquid separator 118, a back pressure controller 120, product gas analysis system shown generally as 122 consisting of a wet test meter 124 and a Ranerax meter 126, and a tempering gas recycle system consisting of recycle gas compressor I28 and recycle gas heater 130.
  • electric heating was provided to the hydrogen humidifier, the vaporizer, the hydrogasifier, and the tempering recycle gas system and all high temperature transfer lines were electrically heated to maintain the process temperature.
  • the gasification system pressure was maintained by the product gas back pressure controller 120.
  • Hydrogasifier temperature was automatically controlled by the hydrogasifier electric trim heaters and the hydrogen feed gas flow was manually set by hand control valve 132 to the desired level as indicated by the rotameter 102.
  • Hydrogen was sparged in beneath the surface of the water 134 maintained in the humidifier and heated by heater 136 to becomesaturated with steam.
  • Oil from tank 108 was pumped by the oil metering pump 110 into electrically heated discharge tubing 138 mixed with the hydrogen and pumped into the electrically heated vaporizer 112.
  • the oil saturated hydrogen vapors exited from the vaporizer 112 through conduit 140 and were introduced into the hydrogasifier 114 through the nozzle 142. Residual unvaporized oil was withdrawn from the vaporizer, as indicated by the level control 144 in order to maintain a minimal level of liquid in the vaporizer, and discarded.
  • the feed inlet tube was jacketed with tempering gas flow, which entered through a pearlite insulated top section 148. Heat leakage from the hydrogasifier 114 was reduced by the pearlite and the heat was largely absorbed by the tempering gas avoiding excessive heating of the feed vapor.
  • the feed gas was injected into the hydrogasifier 114 at the top of the internal draft tube 150 at high velocity. The entering jet induced a large internal recirculation rate to give an approach to complete back-mixing.
  • test hydrogasifier was made of stainless steel reaction vessel approximately 1.16 inches in diameter and 18% inches long.
  • the axial draft tube 150 was a 16-inch length of 20-gauge 96-inch stainless steel tubing.
  • the overall volume of the test hydrogasifier was approximately 19.7 cubic inches.
  • the effluent gases were cooled in a product cooler-condenser 116 and passed through the liquid separator 118 where aromatic condensate liquid and water were collected and periodically drained through conduit 154.
  • the gas from the separator 118 was let down through a back pressure controller and with the rate measured in a wet test meter 124 and specific gravity measured by a Ranerax instrument 126.
  • the test outlined above refers to a successful run feeding 76 weight percent of a medium gravity (32.6 APl) lagomedio crude oil to the gasifier. Approximately 82 weight percent of the oil feed to the gasifier was gasified with 18 weight percent being recovered as aromatic condensate; 76 weight percent vaporization is equivalent to about 1000F TPB cut point but even higher boiling consitutents were sure to be in the gasifier feed. There were no problems experienced in the reactive feed nozzle or vaporizer during this test.
  • the gas recycle hydrogenator reactor system can be used with higher boiling feed stocks including whole crude oil; and it is also believed that residual oil fractions can be used by direct partial vaporization at high pressure into hydrogen of the feed stock. Partial vaporization avoids the deposition and plugging problems that typically result when high boiling distillates or residual petroleum fractions are evaporated to dryness.
  • the residual liquid from the evaporation step with high boiling oil feed stocks provides a suitable feed to a partial oxidation process to supply the hydrogen requirements for hydrogasification.
  • a method of producing a pipeline gas of high heating value from crude oil comprising the steps of:
  • a method according to claim 1 wherein a portion of the product stream is cooled and after removal of residual aromatic and hydrogen sulfide is introduced into the gasification vessel to maintain temperature control in the gasifier vessel.
  • cryogenically separated hydrogen is warmed to ambient temperature and mixed with fresh hydrogen for injection into the gasifier.
  • a method for producing a pipeline gas having a heating value of about 1000 BTU/SCF from crude oil comprising the steps of:
  • a method according to claim 9 wherein the reaction of the naphtha fraction and the methane and ethane effluent with steam is an autothermic catalytic reaction and the naphtha fraction is desulfurized prior to said autothermic catalytic reaction.
  • a method of gasifying a crude oil to produce a gaseous effluent consisting essentially of hydrogen, methane, ethane, hydrogen sulfide and uncondensed aromatics comprising the steps of:
  • a method according to claim 14 wherein a portion of the effluent stream is injected into the hydrogasification vessel to maintain the reaction temperature of said vessel.
  • a method of gasifying a crude oil to produce a gaseous effluent consisting essentially of hydrogen, methane ethane, hydrogen sulfide and uncondensed aromatics comprising the steps of:

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US297012A 1972-10-12 1972-10-12 Method for production of synthetic natural gas from crude oil Expired - Lifetime US3870481A (en)

Priority Applications (13)

Application Number Priority Date Filing Date Title
US297012A US3870481A (en) 1972-10-12 1972-10-12 Method for production of synthetic natural gas from crude oil
CA181,269A CA1001416A (en) 1972-10-12 1973-09-18 Production of synthetic natural gas from crude oil
ZA737505*A ZA737505B (en) 1972-10-12 1973-09-24 Production of synthetic natural gas from crude oil
AU60657/73A AU463986B2 (en) 1972-10-12 1973-09-25 Production of synthetic natural gas from crude oil
GB4512973A GB1453081A (en) 1972-10-12 1973-09-26 Process for producing synthetic natural gas
GB1705876A GB1453082A (en) 1972-10-12 1973-09-26 Method of gasifying crude oil
IN2183/CAL/1973A IN140810B (US06811534-20041102-M00003.png) 1972-10-12 1973-09-26
JP48113167A JPS4970901A (US06811534-20041102-M00003.png) 1972-10-12 1973-10-08
DE19732350666 DE2350666A1 (de) 1972-10-12 1973-10-09 Herstellung von synthetischem naturgas aus rohoel
ES419508A ES419508A1 (es) 1972-10-12 1973-10-09 Un metodo para la produccion de un gas de gran valor calo- rifico.
FR7336265A FR2208967B1 (US06811534-20041102-M00003.png) 1972-10-12 1973-10-10
IT53024/73A IT996290B (it) 1972-10-12 1973-10-10 Metodo per produrre gas metano da petrolio greggio
BE2053138A BE805927A (fr) 1972-10-12 1973-10-11 Procede de fabrication a partir de petrole brut d'un gaz transporte par gazoduc

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Application Number Priority Date Filing Date Title
US297012A US3870481A (en) 1972-10-12 1972-10-12 Method for production of synthetic natural gas from crude oil

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FR2324708A1 (fr) * 1975-09-18 1977-04-15 Air Prod & Chem Procede de gazeification de charges hydrocarbonees
US4025318A (en) * 1975-09-18 1977-05-24 Air Products And Chemicals, Inc. Gasification of hydrocarbon feedstocks
US4046523A (en) * 1974-10-07 1977-09-06 Exxon Research And Engineering Company Synthesis gas production
US4127393A (en) * 1975-01-13 1978-11-28 British Gas Corporation Method and apparatus for vaporizing hydrocarbon based liquids
US4209305A (en) * 1975-12-08 1980-06-24 British Gas Corporation Process for making substitute natural gas
US4300917A (en) * 1978-03-30 1981-11-17 Kraftwerk Union Aktiengesellschaft Method for preventing adhesion or caking of hydrocarbon-containing raw materials
US4345915A (en) * 1977-11-02 1982-08-24 General Electric Company Mixed feed evaporator
US4367077A (en) * 1981-04-20 1983-01-04 Air Products And Chemicals, Inc. Integrated hydrogasification process for topped crude oil
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Cited By (14)

* Cited by examiner, โ€  Cited by third party
Publication number Priority date Publication date Assignee Title
US4046523A (en) * 1974-10-07 1977-09-06 Exxon Research And Engineering Company Synthesis gas production
US4127393A (en) * 1975-01-13 1978-11-28 British Gas Corporation Method and apparatus for vaporizing hydrocarbon based liquids
FR2324708A1 (fr) * 1975-09-18 1977-04-15 Air Prod & Chem Procede de gazeification de charges hydrocarbonees
US4025318A (en) * 1975-09-18 1977-05-24 Air Products And Chemicals, Inc. Gasification of hydrocarbon feedstocks
US4209305A (en) * 1975-12-08 1980-06-24 British Gas Corporation Process for making substitute natural gas
US4345915A (en) * 1977-11-02 1982-08-24 General Electric Company Mixed feed evaporator
US4300917A (en) * 1978-03-30 1981-11-17 Kraftwerk Union Aktiengesellschaft Method for preventing adhesion or caking of hydrocarbon-containing raw materials
US4367077A (en) * 1981-04-20 1983-01-04 Air Products And Chemicals, Inc. Integrated hydrogasification process for topped crude oil
WO2001081723A1 (en) * 2000-04-20 2001-11-01 Scotoil Group Plc Enhanced oil recovery by in situ gasification
US6805194B2 (en) 2000-04-20 2004-10-19 Scotoil Group Plc Gas and oil production
WO2011083297A1 (en) 2010-01-05 2011-07-14 Johnson Matthey Plc Apparatus & process for treating natural gas
US9284236B2 (en) 2010-01-05 2016-03-15 Johnson Matthey Plc Apparatus and process for treating natural gas
US10010858B2 (en) 2010-01-05 2018-07-03 Johnson Matthey Plc Apparatus and process for treating natural gas
WO2024064250A1 (en) * 2022-09-21 2024-03-28 Air Liquide Advance Technologies U.S. Llc Improved system and method for recovery of fuel gas from crude oil purification

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IN140810B (US06811534-20041102-M00003.png) 1976-12-25
BE805927A (fr) 1974-02-01
GB1453082A (en) 1976-10-20
CA1001416A (en) 1976-12-14
AU6065773A (en) 1975-03-27
IT996290B (it) 1975-12-10
JPS4970901A (US06811534-20041102-M00003.png) 1974-07-09
AU463986B2 (en) 1975-08-14
FR2208967A1 (US06811534-20041102-M00003.png) 1974-06-28
ZA737505B (en) 1974-08-28
DE2350666A1 (de) 1974-04-25
ES419508A1 (es) 1976-05-01
GB1453081A (en) 1976-10-20
FR2208967B1 (US06811534-20041102-M00003.png) 1977-05-27

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