US3484496A - Desulphurisation and hydrogenation of aromatic hydrocarbons - Google Patents

Desulphurisation and hydrogenation of aromatic hydrocarbons Download PDF

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US3484496A
US3484496A US583740A US3484496DA US3484496A US 3484496 A US3484496 A US 3484496A US 583740 A US583740 A US 583740A US 3484496D A US3484496D A US 3484496DA US 3484496 A US3484496 A US 3484496A
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hydrogenation
reactor
hydrogen
desulphurisation
sulphur
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John Carruthers
John Winsor
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BP PLC
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BP PLC
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    • 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
    • C10G45/00—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds
    • C10G45/02—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds to eliminate hetero atoms without changing the skeleton of the hydrocarbon involved and without cracking into lower boiling hydrocarbons; Hydrofinishing
    • C10G45/04—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds to eliminate hetero atoms without changing the skeleton of the hydrocarbon involved and without cracking into lower boiling hydrocarbons; Hydrofinishing characterised by the catalyst used
    • C10G45/06—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds to eliminate hetero atoms without changing the skeleton of the hydrocarbon involved and without cracking into lower boiling hydrocarbons; Hydrofinishing characterised by the catalyst used containing nickel or cobalt metal, or compounds thereof

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  • the fraction, together with hydrogen, is contacted at an elevated temperature and pressure with a supported nickel catalyst, the flow of hydrogen being controlled to maintain the equilibrium hydrogen partial pressure greater than the minimum necessary to prevent catalyst deactivation as sulphur is adsorbed by the catalyst but less than that at which up to mol percent hydrogenation of the aromatic hydrocarbons occurs.
  • the desulphurised fraction still under pressure and containing up to 2 p.p.m. wt. sulphur, is hydrogenated over a supported nickel catalyst in two stages, the bulk of the hydrogenation, 90 to 99% wt., occurring in the first stage with cooling so as to control the stage temperature within desired limits, and the hydrogenation completed in the second stage, without coolmg.
  • the present invention is concerned with control of the hydrogen partial pressure during at least the initial stage when the nickel has high hydrogenation activity.
  • a process for the desulphurisation of an aromatic hydrocarbon-conlit) 3,484,496 Patented Dec. 16, 1969 taining fraction containing from 1 to 50 p.p.m. wt. sulphur without appreciable hydrogenation of the aromatic hydrocarbons comprises passing the fraction in the liquid or vapour phase, and in the persence of hydrogen, over supported nickel at an elevated temperature and pressure such that sulphur combines with the nickel but no substantial amount of hydrogen sulphide is produced, the equilibrium hydrogen partial pressure being greater than the minimum necessary to prevent catalyst deactivation but less than that at which up to 10 mol percent hydrogenation of the aromatic hydrocarbons occurs.
  • the invention consists in a process in which an aromatic hydrocarbon-containing fraction containing up to 50 p.p.m. wt. sulphur is desulphurised by the above-mentioned process, to produce a fraction containing up to 2 p.p.m. wt. sulphur, and this fraction is hydrogenated in two stages both using: supported nickel catalysts, in which not less than wt. and not more than 99% wt. of the fraction is hydrogenated in the first hydrogenation stage, with the hydrogenation reaction being substantially completed in the second hydrogenation stage, the temperature of the first hydrogenation stage being controlled by cooling and the second hydrogenation stage being uncooled.
  • the aromatic hydrocarbon-containing fraction need not consist wholly of aromatics, but if it does not the preferred other components are saturated hydrocarbons.
  • the feedstock contains at least wt. of aromatic hydrocarbons.
  • the preferred aromatic hydrocarbon is benzene and subsequent discussion, both of desulphurisation and hydrogenation will be based on the treatment of this substance.
  • the production of cyclohexane by the hydrogenation of benzene is an important industrial process, since cyclohexane is used in the manufacture of nylon, as a solvent and as a reactant in chemical syntheses. Fibre grade cyclohexane, for nylon production, must have a purity of not less than 99.5% and a total aromatic content of less than 500 p.p.m.
  • the present desulphurisation process is capable of dealing with feedstocks containing from 1 to 50 p.p.m. wt., and preferably not more than 10 p.p.m. wt. sulphur in any form, including thiophenic sulphur.
  • Feedstocks containing higher amounts of sulphur may be subjected to any of the known catalytic hydrodesulphurisation processes, using catalysts of poor hydrogenation activity, for example, cobalt and molybdenum oxides on alumina, to reduce the sulphur content to the desired level, or more than one stage of desulphurisation according to the present process may be used and the present desulphurisation process includes such hydrodesulphurisation steps.
  • the hydrogenation stages of the invention may use an aromatic feedstock containing up to 2 p.p.m. wt. sulphur, although the desulphurisation process is capable of producing a product containing less than 1 p.p.m. wt sulphur.
  • Nickel is susceptible to de-activation. by sulphur-containing materials although it has a number of advantages over other substances used for the hydrogenation of aromatic hydrocarbons.
  • the supported nickel catalysts used in the present processes may incorporate any of the known natural or synthetic support materials, such as the refractory oxides of elements of Groups II to V of the Periodic Table, or kieselguhr, pumice, or sepiolite.
  • Sepiolite is the preferred material and the preferred catalyst for both the desulphurisation and hydrogenation processes of the invention is nickel on sepiolite prepared 3 and activated according to the disclosures of British Patent No. 899,652. It is not essential, however, that the same catalyst should be used in the desulphurisation process as in any subsequent hydrogenation process or in both stages of the hydrogenation process.
  • Nickel on sepiolite prepared and activated according to the above-mentioned British patent may contain from 1 to 50% wt. nickel (expressed as elemental nickel) and more particularly from 5 to 25% wt.
  • Such a catalyst has a high nickel surface area and has high activity and selectivity. It is capable of maintaining its hydrogenation activity in the hydrogenation stages of the invention up to a sulphurznickel atomic ratio of 01:1 and its total sulphur capacity is much higher than this.
  • sulphur absorption takes place at least up to 0.75:1 sulphurznickel atomic ratio. Since the sulphur capacit of the supported nickel material is high and is known, it is possible to provide a sufficient amount to give an economic catalyst life. It has been found that a life in excess of one year can be obtained with nickel on sepiolite using a feedstock containing 1.3 p.p.m. wt. thiophenic sulphur.
  • FIGURE 2 It will be seen from FIGURE 2 that there is a significant hydrogen partial pressure at equilibrium at temperatures above about 400 F. (204 C.) even at low levels of conversion. At a hydrogen partial pressure of only 0.5 p.s.i. it has been calculated that there is at least a hundred times as much hydrogen as the minimum required to prevent de-activation of the catalyst surface by polymerisation. Although not shown in FIGURE 2, a hydrogen partial pressure sufficient to prevent de-activation of the nickel surface may be achieved at temperatures as low as 122 F. (50 C.).
  • a conversion of mol percent may be regarded as a satisfactory compromise between the need to avoid excessive hydrogenation and the requirement to provide an adequate hydrogen partial pressure, and reference to FIG- URE 2 will show that such a level of conversion will provide a hydrogen partial pressure of 0.5 p.s.i. at 400 F. A conversion of mol percent will be seen to provide a hydrogen partial pressure slightly greater than this at 400 F.
  • the desulphurisation reaction may be conducted in liquid or vapour phase or in mixed (gas/liquid) phase. Vapour phase operation is preferred. If in the liquid phase, upward flow in the reactor is preferably employed. A fixed or fluidised catalyst bed may be used. In the case of a fluidised catalyst high liquid velocities may cause catalyst to be carried over with the product, in which case a settling tank would be necessary to recover the product.
  • the desulphurisation reaction should preferably be in the vapour phase to allow the hydrogen partial pressure to be at the desired low level.
  • the maximum plant pressure is restricted to the vapour pressure of the benzene at the desired temperature of operation. At 400 F. the plant pressure is thus restricted to above 200 p.s.i.g., and at 450 F. to about 300 p.s.i.g. Operation at fairly high temperatures means that the sulphur capacity and desulphurisation activity of the catalyst are increased and also the equilibrium hydrogen partial pressure is increased.
  • an upper limit of operating temperature is set by the onset of side-reactions, such as cracking, isomerisation, and ring opening. Of these cracking is the most important.
  • the operating temperature can be raised without byproduct formation taking place.
  • Material containing up to 2 p.p.m. wt. sulphur, derived from the desulphurisation process just described, may be hydrogenated over supported nickel in two stages.
  • first of these will be designated the main reactor stage and the second the finishing reactor stage.
  • the material entering the main reactor is preferably in mixed (gas/liquid) phase, and is preferably in the vapour phase on leaving the reactor. It may be in mixed phase or in vapour phase on leaving the reactor, depend ing on the extent of hydrogenation taking place in the reactor and the extent of cooling.
  • the inlet material may possibly be in the vapour phase, but in this case as the outlet temperature is fixed a large amount of recycle cooling would be necessary. Since the bulk of hydrogenation occurs in the first hydrogenation stage i.e. the main reactor, the major part of the heat produced by the hydrogenation reaction is produced in this stage and cooling is therefore necessary to control the stage temperature to within the desired limits. This cooling may be achieved either by liquid recycle or by the use of a cooled tubular reactor.
  • Liquid may be conveniently recycled from the main reactor outlet or from the finishing reactor outlet.
  • the use of liquid recycle means that the linear velocity through the reactor is increased, with consequent enlargement of the reactor to achieve the same contact time.
  • This may be avoided by using a cooled tubular reactor, with the catalyst in the tubes and a cooling agent being passed over them. In this way the temperature rise in the tubes is limited to the required range.
  • a higher average catalyst bed temperature can be attained for a given level of hydrogenation than is possible with an adiabatic reactor.
  • Suitable cooling agents for the cooled tubular reactor are steam, water under pressure, gas, or indeed any substance which is thermally stable within the temperature range of the process.
  • the limiting factor is the rate at which heat can be removed to keep the catalyst at a temperature within the acceptable range. If a cooled tubular reactor is used as the main hydrogenation reactor this stage is preferably in the vapour phase throughout, since otherwise distribution difficulties may occur.
  • the most convenient recycle cooling medium is cyclohexane, and desirably this is, as far as is possible, in the liquid phase at the main reactor inlet, since the heat of vaporisation will assist the cooling effect, and the minimum of cyclohexane to achieve the necessary cooling may be recycled.
  • reaction temperature would have to be less than about 200 C., so that the hydrogen partial pressure would be at a suitable level in relation to the total pressure.
  • reaction temperature would have to be less than about 200 C., so that the hydrogen partial pressure would be at a suitable level in relation to the total pressure.
  • recycle would require recycle at a level which would unduly depress the hydrogen partial pressure.
  • the alternatives of using a higher total pressure or a very large excess of hydrogen would be impracticable in a commercial process.
  • the hydrogenation stages of the invention use an uncooled finishing reactor. From 90 to 99% wt., and preferably about 95% wt. of cyclohexane is produced in the main hydrogenation reactor with the remainder of the conversion taking place in the finishing reactor. At these lower levels of conversion in the main reactor a higher main reactor exit temperature may be employed than if 100% conversion were .attempted in the main reactor, While maintaining a satisfactory hydrogen partial pressure.
  • a higher hydrogen partial pressure is more easily maintained in the finishing reactor than in the main reactor
  • the finishing reactor may operate in mixed (gas/liquid) phase or vapour phase, and its outlet temperature may be, and preferably is, lower than the main reactor outlet temperature, since this is advantageous for high levels of conversion.
  • the hydrogen used in the process of the invention may be commercially pure or it may be a mixed gas derived from a refinery process, such as a steam reformer tail gas, also containing methane, or catalytic reformer off-gas.
  • a steam reformer tail gas also containing methane, or catalytic reformer off-gas.
  • Gases containing two or more carbon atoms per molecule may be used, provided that reaction conditions, and in particular the temperature, are carefully controlled to avoid, in the presence of fresh nickel, a temperature runaway.
  • the gas contains at least 50 mole percent hydrogen, and more suitably 70 to 99 mole percent hydrogen.
  • An advantage of the process is that hydrogen produced by steam reforming of natural gas or naphtha may be used without make-up gas compression.
  • gas may be recycled to the main reactor, and if a mixed gas is used, for example one containing methane, gas may be purged from the recycle gas stream, or not, as desired, or methane may be removed in the liquid leaving the high pressure Temperature, 77 to 662 F. (25 to 350 C.) (preferably 122 to 572 F. (50 to 300 C.))
  • Product recycle ratio 2:1 to 10:1 (preferably 2.5:1 to
  • Finishing reactor Temperature 77 to 662 F. (25 to 350 C.) (preferably 122 to 572 F. (50 to 300 C.))
  • Hydrogen recycle rate to 5000 s.c..f./b. (preferably 500 to 2000 s.c.f./b.
  • the temperature rises occurring in each stage will be within the ranges given, the reactor exit temperatures not exceeding the upper limits of the ranges: set out, and the inlet temperatures being above the lower limits.
  • references made in this specification to main hydrogenation reactors, and finishing reactors, or these terms suffixed by the word stage include the use of one or more reactors in any stage, or the use of one or more reactors containing more than one stage. It is only required that the catalysts of each stage should be physically separate and that independent control of the process parameters should be possible in each stage.
  • FIGURE 1 of the accompanying drawings One embodiment of a combination process in which benzene is desulphurised by the process previously described and the desulphurised benzene is then hydrogenated as described is illustrated in FIGURE 1 of the accompanying drawings as a simplified flowsheet. It should be emphasized that this flowsheet combines desulphurisation and hydrogenation processes although the former is separately viable.
  • benzene enters the system via line 1 and pump 2, and hydrogen or a mixed gas containing hydorgcn enters via line 3 and valve 4.
  • the mixture passes into a desulphurisation reactor 6 and thence into the main reactor 7, make-up hydrogen being supplied via line 3 and another valve 4.
  • main reactor 7 cyclohexane, residual unconverted benzene and excess hydrogen pass into a high pressure separator 8, from which the cyclohexane-benzene mixture is recycled via line and a valve 4 to inlet line 1, if desired, tocontrol the hydrogen partial pressure in the desulphurisation reactor.
  • Gas is recycled via line 9 and cyclohexane-benzene recycled via pump 11 and line 10.
  • a minor proportion of the cyclohexane-benzene passes via line 12 and a valve 4 to the finishing reactor 13 and thence to a high pressure separator 14, from which cyclohexane product passes via line .15 to storage, gas from separator 14 being returned to the main reactor.
  • a mixed gas is used the inerts may be vented at a point on the line between separator 8 and line 9, since the gas leaving separator 8 will have a higher inerts content than that from separator 14.
  • make-up hydrogen may be supplied via line 3 directly to the finishing reactor.
  • typical process conditions might be a common pressure in each reactor of 200 p.s.i.g., respective temperatures in the desulphurisation reactor, main reactor and finishing reactor of 400- 500 R, (204-260" C.), 200-480 F. (93250 C.) and 320-430 F. (160-220 C.) and a space velocity in each reactor of 1 v./v./hr., nickel sepiolite being the catalyst in both the desulphurisation rand hydrogenation processes.
  • the gas used could be steam reformer hydrogen, containing 95% hydrogen and supplied at 250 p.s.i.g., and 95 cyclohexane could be recycled at a recycle ratio to fresh feed of 3.5 :1 to 4.0:1 by weight to control the temperature of the main reactor.
  • a cyclohexane product containing only 1 p.p.m. wt. benzene could be obtained without significant de-acti-vation of the nickel sepiolite.
  • Example I refers to the desulphurisation process of the invention and Examples 2 and 3 refer to the subsequent hydrogenation stages.
  • EXAMPLE 1 The sulphur content of benzene was reduced from 1.3 to 0.2 p.p.m. weight over nickel on sepiolite catalyst under the following conditions:
  • the run was continued under these conditions until 1007 hours on stream, at which time the temperature was lowered to 430 F. At this temperature the sulphur content of the product was 0.3 p.p.m. wt. The run was con tinued under these conditions until 4400 hours on stream, by which time the hot spot had travelled 39 percent down the catalyst bed. If this continued at the same rate the catalyst life would be 11,300 hours or 67 weeks.
  • the purity of the cyclohexane product was 99.78 percent weight.
  • the purity of the final cyclohexane product was within the range 99.86 to 99.92% wt.
  • desulphurising at 450 F. caused by-product formation at a slightly higher level.
  • the cyclohexane product had a sulphur content within the range 0.2 to 0.5 p.p.m. wt. depending on the sulphur content of the feedstock.
  • the rate at which the operating temperature had to be raised to maintain the required degree of hydrogenation is given below:
  • a process for the desulphurisation of a benzenecontaining fraction containing from 1 to 50 p.p.m. wt. sulphur including thiophene sulphur, without appreciable hydrogenation of the aromatic hydrocarbons which comprises passing the fraction in the fluid state, together with hydrogen, over nickel supported on sepiolite at an elevated temperature of from 50 to 250 C. and pressure of from 100 to 500 p.s.i.g. such that sulphur combines with the nickel but no substantial amount of hydrogen sulphide is produced, the equilibrium hydrogen partial pressure being greater than the minimum necessary to prevent catalyst deactivation but less than that at which up to 10 mol percent hydrogenation of the aromatic hydrocarbons occurs.
  • reaction conditions are selected from the following:
  • reaction conditions are selected from the following:
  • reaction conditions for the first and second hydrogenation stages are selected from the following:

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US583740A 1965-10-04 1966-10-03 Desulphurisation and hydrogenation of aromatic hydrocarbons Expired - Lifetime US3484496A (en)

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Cited By (6)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US3767562A (en) * 1971-09-02 1973-10-23 Lummus Co Production of jet fuel
US3992464A (en) * 1974-11-08 1976-11-16 Uop Inc. Hydroprocessing aromatics to make cycloparaffins
US3996304A (en) * 1974-06-19 1976-12-07 Universal Oil Products Company Hydroprocessing of hydrocarbons
US4036734A (en) * 1973-11-02 1977-07-19 Exxon Research And Engineering Company Process for manufacturing naphthenic solvents and low aromatics mineral spirits
WO1996017039A1 (en) * 1994-12-01 1996-06-06 Mobil Oil Corporation Integrated process for the production of reformate having reduced benzene content
US5741414A (en) * 1994-09-02 1998-04-21 Nippon Oil Co., Ltd. Method of manufacturing gas oil containing low amounts of sulfur and aromatic compounds

Citations (14)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US1974057A (en) * 1931-12-11 1934-09-18 Tide Water Oil Company Two stage method of hydrogenation
US2300877A (en) * 1940-08-12 1942-11-03 Phillips Petroleum Co Process for treating hydrocarbons
US2303075A (en) * 1938-11-12 1942-11-24 Phillips Proroleum Company Catalytic hydrogenation process
US2426929A (en) * 1944-07-17 1947-09-02 Shell Dev Hydrogenation of liquid carbonaceous materials
US2459465A (en) * 1945-05-11 1949-01-18 Standard Oil Dev Co Two-stage hydrogenation treatment for hydrocarbon oils
US2464539A (en) * 1945-09-19 1949-03-15 Standard Oil Dev Co Two-stage destructive hydrogenation of petroleum oil
US2755317A (en) * 1952-11-10 1956-07-17 Universal Oil Prod Co Hydrogenation of benzene to cyclohexane
US3147210A (en) * 1962-03-19 1964-09-01 Union Oil Co Two stage hydrogenation process
US3190830A (en) * 1962-03-10 1965-06-22 British Petroleum Co Two stage hydrogenation process
US3202723A (en) * 1961-09-13 1965-08-24 Inst Francais Du Petrole Process for the catalytic hydrogenation of aromatic hydrocarbons
US3215751A (en) * 1961-08-08 1965-11-02 British Petroleum Co Isomerisation of olefins
US3222274A (en) * 1963-01-02 1965-12-07 Socony Mobil Oil Co Inc Process for producing high energy jet fuels
US3254134A (en) * 1965-04-05 1966-05-31 Texaco Inc Plural stage hydrogenation of aromatics
US3341613A (en) * 1964-10-30 1967-09-12 Phillips Petroleum Co Method for production of cyclohexane by the hydrogenation of benzene

Patent Citations (14)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US1974057A (en) * 1931-12-11 1934-09-18 Tide Water Oil Company Two stage method of hydrogenation
US2303075A (en) * 1938-11-12 1942-11-24 Phillips Proroleum Company Catalytic hydrogenation process
US2300877A (en) * 1940-08-12 1942-11-03 Phillips Petroleum Co Process for treating hydrocarbons
US2426929A (en) * 1944-07-17 1947-09-02 Shell Dev Hydrogenation of liquid carbonaceous materials
US2459465A (en) * 1945-05-11 1949-01-18 Standard Oil Dev Co Two-stage hydrogenation treatment for hydrocarbon oils
US2464539A (en) * 1945-09-19 1949-03-15 Standard Oil Dev Co Two-stage destructive hydrogenation of petroleum oil
US2755317A (en) * 1952-11-10 1956-07-17 Universal Oil Prod Co Hydrogenation of benzene to cyclohexane
US3215751A (en) * 1961-08-08 1965-11-02 British Petroleum Co Isomerisation of olefins
US3202723A (en) * 1961-09-13 1965-08-24 Inst Francais Du Petrole Process for the catalytic hydrogenation of aromatic hydrocarbons
US3190830A (en) * 1962-03-10 1965-06-22 British Petroleum Co Two stage hydrogenation process
US3147210A (en) * 1962-03-19 1964-09-01 Union Oil Co Two stage hydrogenation process
US3222274A (en) * 1963-01-02 1965-12-07 Socony Mobil Oil Co Inc Process for producing high energy jet fuels
US3341613A (en) * 1964-10-30 1967-09-12 Phillips Petroleum Co Method for production of cyclohexane by the hydrogenation of benzene
US3254134A (en) * 1965-04-05 1966-05-31 Texaco Inc Plural stage hydrogenation of aromatics

Cited By (6)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US3767562A (en) * 1971-09-02 1973-10-23 Lummus Co Production of jet fuel
US4036734A (en) * 1973-11-02 1977-07-19 Exxon Research And Engineering Company Process for manufacturing naphthenic solvents and low aromatics mineral spirits
US3996304A (en) * 1974-06-19 1976-12-07 Universal Oil Products Company Hydroprocessing of hydrocarbons
US3992464A (en) * 1974-11-08 1976-11-16 Uop Inc. Hydroprocessing aromatics to make cycloparaffins
US5741414A (en) * 1994-09-02 1998-04-21 Nippon Oil Co., Ltd. Method of manufacturing gas oil containing low amounts of sulfur and aromatic compounds
WO1996017039A1 (en) * 1994-12-01 1996-06-06 Mobil Oil Corporation Integrated process for the production of reformate having reduced benzene content

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FR1498696A (fr) 1967-10-20
NO120151B (en)) 1970-09-07
NL6613968A (en)) 1967-04-05
GB1141809A (en) 1969-02-05
BE687782A (en)) 1967-04-04
DE1568189C3 (de) 1975-01-02
DE1568189B2 (en)) 1974-05-22
ES332400A1 (es) 1967-08-01
CH508563A (de) 1971-06-15
DE1568189A1 (de) 1970-02-05
AT267026B (de) 1968-12-10
IL26615A (en) 1970-07-19
SE346556B (en)) 1972-07-10

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