US2672433A - Catalytic desulfurization of petroleum hydrocarbons - Google Patents

Catalytic desulfurization of petroleum hydrocarbons Download PDF

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US2672433A
US2672433A US213176A US21317651A US2672433A US 2672433 A US2672433 A US 2672433A US 213176 A US213176 A US 213176A US 21317651 A US21317651 A US 21317651A US 2672433 A US2672433 A US 2672433A
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hydrogen
sulphur
zone
desulphurisation
residue
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Porter Frederick Willi Bertram
Rowland John
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Anglo Iranian Oil Co Ltd
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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
    • C10G65/00—Treatment of hydrocarbon oils by two or more hydrotreatment processes only
    • C10G65/14—Treatment of hydrocarbon oils by two or more hydrotreatment processes only plural parallel stages only
    • C10G65/16—Treatment of hydrocarbon oils by two or more hydrotreatment processes only plural parallel stages only including only refining steps

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  • This invention relates to the hydrocatalytic desulphurisation of petroleum hydrocarbons.
  • Desulphurisation processes are also known in which a substance capable of supplying hydrogen under the conditions of the desulphurisation reaction is added to the feedstock to be desulpnurised, but this is only a special case of hydrofining and suffers from the disadvantage of the cost of supplying the hydrogen donor and from the additional disadvantage that the through-put of the feedstock is reduced .by the addition of the considerable quantity of hydrogen donor necessary to supply suflicient hydrogen.
  • the preferred catalyst for use in the autofining process is of the so-called cobalt molybdate type which comprises mixtures of the oxides of cobalt and molybdenum, or chemical compounds of cobalt, molybdenum and oxygen, or mixtures of one or both of said oxides with said compounds either alone or incorporated with a support. It has been found that there is a period at the commencement of each-run during which the gas make is negligible or non-existent and the desulphurisation is not at its maximum.
  • the regeneration of the used catalyst may be carried out by burning off the carbon and sulphur deposits with either a nitrogen/air or steam/air mixture.
  • the following conditions may be employed for regeneration.
  • the autofining process may be carried out in a static bed reactor or by the use of the moving bed or fluid catalyst techniques.
  • an autofin n process is meant a process m whi h Pe oleum, e d o i passed in v ncur, Q m.-;. n.
  • the 'autofin'i'ng zone may fibz-jtestisieient,ts sdppiy all the'hydrogen 'resuites tithe mal easanc zone admin in any case represent a'considerable saving in hydrogen.
  • he independent-hydrogenation zone may 'coniiis't'of a-hyd'rofining-zone for the treatment of Heavier petroleum -feedstocks such as wa x -dis'- r tillate, crude on 'and reducederude, that cannot be satisfactorily autofined.
  • tliehydrogenation zone may b'efor carrying out a well known hydrogenation process; such as the conversion of 'dii's'dbutylene to 'is'ooetane.
  • --a-crude -oil is fractionated for the recovery of a -light fraction, which isfautofined and a. residue, the excess gas make trom the autofining stage being used to hydrofinethe residueor a heavy fraction thereof.
  • the naphtha-kerosine blend is fed by a pump l to the preheater 2 after admixture with the recycle gas.
  • the hot vapours pass downwards through the autofining reactor 4 and after passing through the cooler or heat exchanger 5 are separated into two phases in the vessel 6.
  • the liquid phase which contains the desulphurised kerosine-naphtha blend is passed via line I to a stabiliser and distillation unit for separation of the fractions as desired.
  • the gas phase which contains upwards of per cent mol. hydrogen is split into two streams, one being recycled through line 3 and the booster 3 to maintain the conditions in the reactor 4 while the other, the excess gas, passes through a control valve 9.
  • the wax-distillate fraction is fed by a pump to into admixture with the excess gas from the autofining system and passes to the preheater H via line it.
  • the preheated products pass downward over the catalyst bed 12 through a pressure control valve 13 and cooler id to the separator IS.
  • the liquid phase is run ofi via line I! for stabilisation for removal of hydrogen sulphide and the gases vented via line 18.
  • the reactor 4 and its accompanying system are operated at a, pressure of 220 pounds per square inch pressure and the excess gas fed into the wax-distillate stream at +200 pounds per square inch.
  • the wax distillate section therefore operates as a once through hydrofining process i. 6. with no recycle, and pressure is supplied from the first section so that an economic hydrogen partial pressureis maintained in the second section. If necessary, the relative proportion of hydrogen in the excess gas can be increased by scrubbing out the hydrocarbons but this is not essential, neither is it essential to remove hydro gen sulphide from the recycle or excess gas.
  • This method of operation is a'specific example of this method of operation:
  • War-distillate desulphurisatiqn The method illustrated in Figure 2 is essentially the same as that illustrated in Figure l, the only difierence being the provision of a compressor l9 between the two sections to enable the wax-distillate desulphurisation to be carried out at a higher pressure and thereby increase the sulphur removal.
  • the amount of excess gas from the autofining section can be controlled, within limits, by adjustment of the space velocity. The following is a specific example of this method of operation:
  • Wax-distillate desulphurz'sation Korean wax-distil- Feedstock late cut 72%- 82% v o l o n crude.
  • Catalyst Mixed cobalt and molybdenum oxides on alumina Catalyst Mixed cobalt and molybdenum oxides on alumina.
  • crude petroleum stabilised to C. is fed via preheater 20 to the fractionating tower -2 I where a naphtha-kerosine cut boiling between 20 and 250 C. is taken overhead through condenser 22 to the reflux drum 23.
  • Some of'the naphtha-kerosine is returned to the tower as reflux by the reflux pump 24, the remainder being fed by the feed pump 25 to the preheater 26.
  • the naphtha-kerosine, together with the recycle gas, is brought up to reaction temperature and passes into the reactor 21.
  • the autofined product from the reactor passes out through a waste heat boiler 28 to a knock-out tower 29.
  • the pressure on the waste heat boiler is controlled to give an inlettemperature into the knock-out tower such that the kerosine is condensed whilst the naphtha and recycle gas pass off the top of the knock-out tower through a cooler 30 to a separator 3
  • the recycle gas separates from the liquid naphtha and is recycled by means of the gas booster 32 to the preheater inlet. Further control of the knock-out. tower is effected by returning some of the cold. naphtha from the separator as reflux by the.- pump 33.
  • the kerosine passes from the base of the knock-out tower through a cooler 34 to stor-- age.
  • the naphtha passes from the separator to a conventional stabiliser (not shown) where thedissolved hydrogen sulphide is removed.
  • The.- excess gas over and above that required to main-- tain the system under pressure, is released through the pressure control valve 35 to the compressor 36. Here, it is compressed to the required operating pressure and passes via the flow controller 3'! into the preheater inlet line, where it meets the residue from the fractionator 2
  • the gas-residue mixture is brought up to the reaction temperature in the preheater 38 and passes into the reactor 39.
  • the desulphurised product and the gas pass out via a pressure control valve 40, which can be used to control the preheater and reactor pressure or which can, if required, be bypassed.
  • the hydrocarbon-hydrogen stream passes through a cooler 4
  • the gas can either be vented via the pressure controller 43 or be recycled by the recycle booster 44 to the preheater inlet.
  • the desulphurised residue is' passed from the separator 42 via a stabiliser not shown) to remove dissolved hydrogen sulphide to storage. If required, fresh hydrogen from an external source ma beinjected into the suction of, the compressor 36 through line 45.
  • control valves 40, 43 and 31 and the gas recycle booster 44 depends on whether it is desired to operate with or without a gas recycle system and/or hydrogen added from an external source. This may be summarised as follows:
  • a process for the recovery of sulphur free products from naphthene-containing crude petroleum by the hydrocatalytic.desulphurisation process without the need for an extraneous source of hydrogen which comprises fractionating the crude petroleum for therecovery of a selected naphthene-containing low-boiling fraction and a.
  • said temperature and pressure being correlated so that organically combined sulphur in said fraction is converted into hydrogen sulphide and hydrogen is produced by dehydrogenation of naphthenes contained in said low-boiling fraction in an amount in excess or that required to convert organically combined sulphur contained in said low-boiling fraction into hydrogen sulphide and maintain the necessary partial pressure of hydrogen in said zone, separating hydrogen sulphide and a lwdrogen-rich gas mixture from the products of said first desulphurisation zone, recycling a portion of said hydrogen-rich gas mixture to said first desulphurisation zone to constitute the Whole of the hydrogen supplied to said zone, the hydrogen recycle rate being sufiicient to maintain the necessary partial pressure of hydrogen in said zone therein, passing another portion of said hydrogen-rich gas mixture in admixture with said residue to a second desulphurisation zone wherein said residue is contacted with a sulphur-resistant hydrogenation catalyst and hydrogen of the mixture at a temperature and pressure suitable for hydrofining of said residue whereby organic
  • a process in accordance with claim 1 wherein the recovered low-boiling fraction comprises the naphtha and kerosene fractions.
  • the low-boiling fraction comprises the naphtha and kerosene fractions and the dehydrogenationhydrogenation catalyst is a cobalt molybdate type.

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  • Chemical & Material Sciences (AREA)
  • Oil, Petroleum & Natural Gas (AREA)
  • Engineering & Computer Science (AREA)
  • Chemical Kinetics & Catalysis (AREA)
  • General Chemical & Material Sciences (AREA)
  • Organic Chemistry (AREA)
  • Production Of Liquid Hydrocarbon Mixture For Refining Petroleum (AREA)

Description

March 16, 1954 F. W. B. PORTER ET AL CATALYTIC DESULFURIZATION OF PETROLEUM HYDROCARBONS Filed Feb. 28 1951 Jam:
2 Sheets-Sheet l FEEDER/CK WALL/9M BER RQY PURDY AZOETHCOT'T JOH/V E0 W4 AN TR/IM 1 0/? 72' R March 16, 1954 F. w. B. PORTER ET AL 2,672,433
CATALYTIC DESULFURIZATION OF PETROLEUM HYDROCARBONS Filed Feb. 28, 1951 2 Sheets-Sheet 2 FR'DER/C K W/LA/HM 5E7EAM PQPTEE P0-Y PUED Y NOE THCO T T 20% E0 WL A N0 Patented Mar. 16, 1954 UNIT-EDMSTATES 2,672,433 PATENT OFF-ICE Application February 28, 1951, Serial No. 213,176
Claims priority; application Great Britain March 15, 1950 I Claims. (Cl. 196-28) 1 p This invention relates to the hydrocatalytic desulphurisation of petroleum hydrocarbons.
Among the processes which are known for the removal of organically combined sulphur from petroleum hydrocarbons is the so called hydrofining process in which the petroleum hydrocarbons to be desulphurised are, passed in admixture with hydrogen over a sulphur-resistant hydrogenation catalyst at elevated temperature and pressure whereby the organically combined sulphur is converted into hydrogen sulphidewhich may readily be removed from the treated hydrocarbons the properties of which are not otherwise appreciably affected. This process is technically effective for the removal of organically combined sulphur from petroleum hydrocarbons but is not commercially attractive in View of the cost of supplying the considerable quantity of hydrogen l consumed in the process' It was also known that the hydroforming process produced hydrogen and at the same time effected a considerable degree of desulphurisation, but by its very nature this process is not applicable in cases where it is desired to desulphurise feedstocks, such as gas oils and aromatic extracts, without appreciably affecting the properties of the feedstock other than changes consequent upon the removal of the organically combined sulphur as hydrogen sulphide. Desulphurisation processes are also known in which a substance capable of supplying hydrogen under the conditions of the desulphurisation reaction is added to the feedstock to be desulpnurised, but this is only a special case of hydrofining and suffers from the disadvantage of the cost of supplying the hydrogen donor and from the additional disadvantage that the through-put of the feedstock is reduced .by the addition of the considerable quantity of hydrogen donor necessary to supply suflicient hydrogen.
It was then discovered that by careful control of the temperature and pressure and by using a sufficiently active dehydrogenation-hydrogenation catalyst, it was possible to derive suificient hydrogen by dehydrogenation of naphthenes contained in the feedstock to enable suificient of the organically combined sulphur in the feedstock to be converted into hydrogen sulphide, under the same temperature and pressure conditions, to constitute an effective desulphurisation process. This process has been autofining and it has the considerable advantage as compared with hydrofiningthat all the hydrogen required for desulphurisation is derived, from the feedstock itself. It was surprising to find that conditions existed under which the two reactions of dehydrogenation of naphthenes and. hydrogenation of organic sulphur compounds. could 2v 7 proceed simultaneously .to the extent necessary to provide a satisfactory catalytic desulphurisation process. The process was found to be applicable to a wide variety of feedstocks ranging from naphthas to waxdistillates under the following set of conditions: 1
Pressure p. s. i. ga
Temperature F-.. 750-800 Space velocity ..v./v./hr 1.0-5.0 Recycle rate C. F./B 2000-4000 The preferred catalyst for use in the autofining process is of the so-called cobalt molybdate type which comprises mixtures of the oxides of cobalt and molybdenum, or chemical compounds of cobalt, molybdenum and oxygen, or mixtures of one or both of said oxides with said compounds either alone or incorporated with a support. It has been found that there is a period at the commencement of each-run during which the gas make is negligible or non-existent and the desulphurisation is not at its maximum. This low dehydrogenating activity of the catalyst is more noticeable at low feedstock velocities and with the heavier feedstocks. It was discovered that the non-activity of the catalyst during the early hours on stream could be largely overcome by subjecting the catalyst tothe action of hydrogen sulphide or hydrogen sulphide-containing gas prior to its use in the autofining process. 1
The regeneration of the used catalyst may be carried out by burning off the carbon and sulphur deposits with either a nitrogen/air or steam/air mixture. The following conditions may be employed for regeneration.
Inlet temperature F 800-850 Steam or nitrogen flow v./v./hr 700-900 Air flow v./v./hr 50-60 Inlet oxygen percent mol 1 to 1.5
to maintain the predetermined pressure. In this case, there is a continuous make of hydrogen; indicating that the hydrogen produced inthe dehydrogenation reaction is not being fully utilized in the desulphurisationreaction. Animproved method of operation was therefore d6V1r-;-
w i in w ch iiiehy s n-cqn ai e aseous fraction is i'ebycl'efd the reaction zone and the pressure thereiii allowed to rise to an equilibrium pressure at which the hydrogen evolved equals the hydrogen consumed. This method of opera tion results in a greater degree of desuign ssation and increased on-stream hours for a product of given sulphur content.
The autofining process may be carried out in a static bed reactor or by the use of the moving bed or fluid catalyst techniques. M I
In the autofining of the light straight-runfe'ed stocks, such as naphthas and kerosines', there is a considerable excess gas make which contains upwards of 80 per cent mol. hydrogen. lie; sulphurisation of these light feedstocks is of the I distiuates. crude hi1 and deduced crude, it is necessary to provide large quantities of hydrogen q e nal e tmi -;A s. r s e. re ent inv nto h e c ss sa z b al. i q Whinin l t p t ole heada he 9. 49 1 na htha. and ere i es is used to supply hydrogen to an independent hyosena o zo ah By an autofin n processis meant a process m whi h Pe oleum, e d o i passed in v ncur, Q m.-;. n. dm x ure w h h en v r aeata ystyh h .semh nes a t f the d hyereeenei h 9 ...eanhihnes a oma w t as v ty t e h ims nai n 1 wee e s l- B14131. seaweeds. n which i n t poisoned as asata y tbmh e aq sulp r c m n s. under co emns-9i. em e a re an s r w qhre cee relai dcsp. a de r e a tswhthezia. cee i e iaihe .iee q is seede -.t .-.ae;e. e t Jest su aei a ly in, ex 9.? that sed iq uce. suf c t h d w. as v r i sei ie lly seawe s h in th smash h dr n. sulph d a d. t m a the pre s e the e ct o zone the ydr e v i nsflnhife I es s f o he treated racemes and the hydrogen recycled to the reaction zone as thsble source of hydrogen e excess gas the 'autofin'i'ng zone may fibz-jtestisieient,ts sdppiy all the'hydrogen 'resuites tithe mal easanc zone admin in any case represent a'considerable saving in hydrogen.
he independent-hydrogenation zonemay 'coniiis't'of a-hyd'rofining-zone for the treatment of Heavier petroleum -feedstocks such as wa x -dis'- r tillate, crude on 'and reducederude, that cannot be satisfactorily autofined. On the other hand, tliehydrogenation zone may b'efor carrying out a well known hydrogenation process; such as the conversion of 'dii's'dbutylene to 'is'ooetane.
According to a preferred method of carrying the-invention into effect, --a-crude -oil is fractionated for the recovery of a -light fraction, which isfautofined and a. residue, the excess gas make trom the autofining stage being used to hydrofinethe residueor a heavy fraction thereof.
The invention will now be described with refereiice l to the accompanyin diagrams, wherein Figures- 1:3 illiistrate diiierent methods at carryme the invention into efiect.
='Irl the methods illustrated in Figures land 2, a haphtha heiosine fraction is autoiine'd and the excess gas s dsesto-h'ydrofi'ne a wax distillate frames. The relative proportions by volume r napnt aykerbsipe and tiig its}; distillate out are approximately 2:1.5 and 2:1 for two representative Iranian crudes, although this may vary considerably according to the length of the wax distillate cut. Thus, if the excess gas make is 50 cubic feet per barrel when autofining the naphthawkerosine blend, then there are available 75-100 cubic feet per barrel for the desulphurisation of the wax-distillate or other heavy fraction. Two methods of utilising this excess hydrogen-containing gas for the desulphurisation of wax distillates are illustrated in Figures 1 and 2.
Referring first to Figure l, the naphtha-kerosine blend is fed by a pump l to the preheater 2 after admixture with the recycle gas. The hot vapours pass downwards through the autofining reactor 4 and after passing through the cooler or heat exchanger 5 are separated into two phases in the vessel 6. The liquid phase which contains the desulphurised kerosine-naphtha blend is passed via line I to a stabiliser and distillation unit for separation of the fractions as desired. The gas phase which contains upwards of per cent mol. hydrogen is split into two streams, one being recycled through line 3 and the booster 3 to maintain the conditions in the reactor 4 while the other, the excess gas, passes through a control valve 9. The wax-distillate fraction is fed by a pump to into admixture with the excess gas from the autofining system and passes to the preheater H via line it. The preheated products pass downward over the catalyst bed 12 through a pressure control valve 13 and cooler id to the separator IS. The liquid phase is run ofi via line I! for stabilisation for removal of hydrogen sulphide and the gases vented via line 18. v
The reactor 4 and its accompanying system are operated at a, pressure of 220 pounds per square inch pressure and the excess gas fed into the wax-distillate stream at +200 pounds per square inch. The wax distillate section therefore operates as a once through hydrofining process i. 6. with no recycle, and pressure is supplied from the first section so that an economic hydrogen partial pressureis maintained in the second section. If necessary, the relative proportion of hydrogen in the excess gas can be increased by scrubbing out the hydrocarbons but this is not essential, neither is it essential to remove hydro gen sulphide from the recycle or excess gas. The following is a'specific example of this method of operation:
EXAMPLE 1 Au'tofining section Feedstock .Iranian naphthakerosine fraction 23-48% vol. on
V crude. Catalyst -s Mixed cobalt and molybdenum oxides on alumina.
War-distillate desulphurisatiqn The method illustrated in Figure 2 is essentially the same as that illustrated in Figure l, the only difierence being the provision of a compressor l9 between the two sections to enable the wax-distillate desulphurisation to be carried out at a higher pressure and thereby increase the sulphur removal. The amount of excess gas from the autofining section can be controlled, within limits, by adjustment of the space velocity. The following is a specific example of this method of operation:
EXAMPLE 2 Autofining section Feedstock Iranian naphthakerosine fraction 848% vol. on
crude. Catalyst Mixed cobalt and molybdenum oxides on alumina. Pressure 150 p. s. i. ga. Temperature 780 F. Space velocity 1.0 v./v./hr. Gas recycle rate 2500 S. C. F./B. Gas hydrogen content 85-90% mol. Excess gas 50 S. C. F./B.
Sulphur content of feed 0.25% wt. Sulphur content of product .002% wt. Sulphur removal 99%.
Wax-distillate desulphurz'sation Iranian wax-distil- Feedstock late cut 72%- 82% v o l o n crude.
Catalyst Mixed cobalt and molybdenum oxides on alumina.
Duration of run 50 hours.
Pressur 500 p. s. i. ga.
Temperature 750 F.
Space velocity 1.0 v./v./hr.
Inlet gas rate 200 S. C. F./B.
Sulphur content of feed-.." 1.5% wt. Sulphur content of product- 0.3% wt. Sulphur removal 80%.
It will be noted that to increase the excess gas make in the autofining section, the space velocity and pressure have been decreased.
Referring now to Figure 3, crude petroleum stabilised to C. is fed via preheater 20 to the fractionating tower -2 I where a naphtha-kerosine cut boiling between 20 and 250 C. is taken overhead through condenser 22 to the reflux drum 23.
Some of'the naphtha-kerosine is returned to the tower as reflux by the reflux pump 24, the remainder being fed by the feed pump 25 to the preheater 26. The naphtha-kerosine, together with the recycle gas, is brought up to reaction temperature and passes into the reactor 21. The autofined product from the reactor passes out through a waste heat boiler 28 to a knock-out tower 29. The pressure on the waste heat boiler is controlled to give an inlettemperature into the knock-out tower such that the kerosine is condensed whilst the naphtha and recycle gas pass off the top of the knock-out tower through a cooler 30 to a separator 3|. There the recycle gas separates from the liquid naphtha and is recycled by means of the gas booster 32 to the preheater inlet. Further control of the knock-out. tower is effected by returning some of the cold. naphtha from the separator as reflux by the.- pump 33. The kerosine passes from the base of the knock-out tower through a cooler 34 to stor-- age. The naphtha passes from the separator to a conventional stabiliser (not shown) where thedissolved hydrogen sulphide is removed. The.- excess gas, over and above that required to main-- tain the system under pressure, is released through the pressure control valve 35 to the compressor 36. Here, it is compressed to the required operating pressure and passes via the flow controller 3'! into the preheater inlet line, where it meets the residue from the fractionator 2| and passes into the preheater 38. The gas-residue mixture is brought up to the reaction temperature in the preheater 38 and passes into the reactor 39.
- The desulphurised product and the gas pass out via a pressure control valve 40, which can be used to control the preheater and reactor pressure or which can, if required, be bypassed. The hydrocarbon-hydrogen stream passes through a cooler 4| to a separator 42, where the gas separates from the liquefied product. The gas can either be vented via the pressure controller 43 or be recycled by the recycle booster 44 to the preheater inlet.
The desulphurised residue is' passed from the separator 42 via a stabiliser not shown) to remove dissolved hydrogen sulphide to storage. If required, fresh hydrogen from an external source ma beinjected into the suction of, the compressor 36 through line 45. The use of control valves 40, 43 and 31 and the gas recycle booster 44 depends on whether it is desired to operate with or without a gas recycle system and/or hydrogen added from an external source. This may be summarised as follows:
Case 1.Recycle system without external hy drogen.--Booster 44 and pressure controller 43 in use. Valves 40 and 31 lay-passed.
Case 2.-Recycle system with external hydrogen.Booster 44, pressure controller 43 and flow controller 31 in use. Valve 40 by-passed.
Case 3.Once-through system without external hydrogen-Valve 40 in use, valves 43 and 31 by-passed and booster 44 shut down.
The following results were obtained for stabilised 20 C. Iranian crude:
Yield distillate (naphtha-kero.) BJB. stable crude=.43. Yield residue 250 C.) BJB. stable crude-=57.
Conditions for naphtha-kerdautofining Pressure, p. s. 1. ga -e 100 Reactor temp., F 800 Space velocity, v./v./hr 1.0 Recycle rate, S. C. FJB e000 Total sulphur in feed, percent wt; .089
Total sulphur in naphtha product- 001 Total sulphur in kerosine' .001 Sulphur removal, percent 98.8
Excess gas (approx. 80 mol. percent Hz) 150 S. C. F./B. naphtha-kero sine =64.5 S. C. F./B. stable crude. :113 S. C. F./B. residue.
Conditions for residue desulphurz'sation Case 3 Pressure, p. s. i. 000 500. Reactor Temp, r 750 750. Space Velocity, v./v./h 1.0 1.0. Recycle Rate, S. O. F. l, 000 once through. External H2, S. G. FJB. nil 8 nil Total Sulphur in Feed, percent V6.1 1.81 1. 81 1.8 1. Total Sulphur in Product.. 1.09 0. 36 0.9. Sulphur Removal, percent 40 80 50.
Inspection data. on residue before and after desulphurz'sation (conditions'asfor Case 2) RESIDUE (BEFORE) The product distribution in Casesl and 3 will be similar to Case 2, except that less lower boiling material will be formed.
Among the advantages of-the process according to the invention are the production of a naphtha-kerosine'fraction of low sulphur content, and of a reduced crude of low sulphur content, of lowered viscosity and with an increased content of lower boiling material.
We-claim:
1. A process for the recovery of sulphur free products from naphthene-containing crude petroleum by the hydrocatalytic.desulphurisation process without the need for an extraneous source of hydrogen, which comprises fractionating the crude petroleum for therecovery of a selected naphthene-containing low-boiling fraction and a. residue, passing said low-boiling fraction in admixture with hydrogen derived solely from said fraction to a first desulphurisation zone wherein it is contacted with a sulphur-resistant dehydrogenation-hydrogenation catalyst which is immuneto sulphur poisoning and combines activity for the dehydrogenation of naphthenes to aromatics with activity for the hydrogenation of organically combined sulphur in said fraction to hydrogen sulphide at a selected temperature within the range 750-800 F. and a selected pressure within the range -250 p. s. 1. ga., said temperature and pressure being correlated so that organically combined sulphur in said fraction is converted into hydrogen sulphide and hydrogen is produced by dehydrogenation of naphthenes contained in said low-boiling fraction in an amount in excess or that required to convert organically combined sulphur contained in said low-boiling fraction into hydrogen sulphide and maintain the necessary partial pressure of hydrogen in said zone, separating hydrogen sulphide and a lwdrogen-rich gas mixture from the products of said first desulphurisation zone, recycling a portion of said hydrogen-rich gas mixture to said first desulphurisation zone to constitute the Whole of the hydrogen supplied to said zone, the hydrogen recycle rate being sufiicient to maintain the necessary partial pressure of hydrogen in said zone therein, passing another portion of said hydrogen-rich gas mixture in admixture with said residue to a second desulphurisation zone wherein said residue is contacted with a sulphur-resistant hydrogenation catalyst and hydrogen of the mixture at a temperature and pressure suitable for hydrofining of said residue whereby organically combined sulphur contained in said residue is converted into hydrogen sulphide, said hydrogen-rich gas mixture constituting the sole source of hydrogen supplied to said hydrofining stage, separating hydrogen sulphide from the products of said second desulphurisation zone, and recovering a substantially sulphur-free product from both said desulphurisation zones.
2. A process in accordance with claim 1 in which a hydrogen-rich gas mixture is obtained from the second desulphurisation zone and is recycled to the second desulphurisation zone.
3. A process in accordance with claim 1 wherein the recovered low-boiling fraction comprises the naphtha and kerosene fractions.
4. A process in accordance with claim 1 wherein the sulphur-resistant dehydrogenation-hydrogenation catalyst is a cobalt molybdate type.
5. A process inaccordance with claim 1 wherein the low-boiling fraction comprises the naphtha and kerosene fractions and the dehydrogenationhydrogenation catalyst is a cobalt molybdate type. FREDERICK WILLIAM BER'I'RAM PORTER. ROY PURDY NORTHCOTT.
.JOI-IZN ROWLAND.
References Cited in the file of this patent UNITED STATES PATENTS Number Name Date 2,417,308 Lee Mar. 11, 1947 2,486,361 Nahin et al Oct. 25, 1949 2,567,252 Strang Sept. 11, 1951 2,573,726 Porter et a1 NOV. 6, 1951

Claims (1)

1. A PROCESS FOR THE RECOVERY OF SULPHUR-FREE PRODUCTS FROM NAPHTHENE-CONTAINING CRUDE PETROLEUM BY THE HYDROCATALYTIC DESULPHURISATION PROCESS WITHOUT THE NEED FOR AN EXTRANEOUS SOURCE OF HYDROGEN, WHICH COMPRISES FRACTIONATING THE CRUDE PETROLEUM FOR THE RECOVERY OF A SELECTED NAPHTHENE-CONTAINING LOW-BOILING FRACTION AND A RESIDUE, PASSING SAID LOW-BOILING FRACTION IN ADMIXTURE WITH HYDROGEN DERIVED SOLELY FROM SAID FRACTION TO A FIRST DESULPHURISATION ZONE WHEREIN IT IS CONTACTED WITH A SULPHUR-RESISTANT DEHYDROGENATION-HYDROGENATION CATALYST WHICH IS IMMUNE TO SULPHUR POISONING AND COMBINES ACTIVITY FOR THE DEHYDROGENATION OF NAPHTHENES TO AROMATICS WITH ACTIVITY FOR THE HYDROGENATION OF ORGANICALLY COMBINED SULPHUR IN SAID FRACTION TO HYDROGEN SULPHIDE AT A SELECTED TEMPERATURE WITHIN THE RANGE 750-800* F. AND A SELECTED PRESSURE WITHIN THE RANGE 50-250 P. S. I. GA., SAID TEMPERATURE AND PRESSURE BEING CORRELATED SO THAT ORGANICALLY COMBINED SULPHUR IN SAID FRACTION IS CONVERTED INTO HYDROGEN SULPHIDE AND HYDROGEN IS PRODUCED BY DEHYDROGENATION OF NAPHTHENES CONTAINED IN SAID LOW-BOILING FRACTION IN AN AMOUNT IN EXCESS OF THAT REQUIRED TO CONVERT ORGANICALLY COMBINED SULPHUR CONTAINED IN SAID LOW-BOILING FRACTION INTO HYDROGEN SULPHIDE AND MAINTAIN THE NECESSARY PARTIAL PRESSURE OF HYDROGEN IN SAID ZONE, SEPARATING HYDROGEN SULPHIDE AND A HYDROGEN-RICH GAS MIXTURE FROM THE PRODUCTS OF SAID FIRST DESULPHURISATION ZONE, RECYCLING A PORTION OF SAID HYDROGEN-RICH GAS MIXTURE TO SAID FIRST DESULPHURISATION ZONE TO CONSTITUTE THE WHOLE OF THE HYDROGEN SUPPLIED TO SAID ZONE, THE HYDROGEN RECYCLE RATE BEING SUFFICIENT TO MAINTAIN THE NECESSARY PARTIAL PRESSURE OF HYDROGEN IN SAID ZONE THEREIN, PASSING ANOTHER PORTION OF SAID HYDROGEN-RICH GAS MIXTURE IN ADMIXTURE WITH SAID RESIDUE TO A SECOND DESULPHURISATION ZONE WHEREIN SAID RESIDUE IS CONTACTED WITH A SULPHUR-RESISTANT HYDROGENATION CATALYST AND HYDROGEN OF THE MIXTURE AT A TEMPERATURE AND PRESSURE SUITABLE FOR HYDROFINING OF SAID RESIDUE WHEREBY ORGANICALLY COMBINED SULPHUR CONTAINED IN SAID RESIDUE IS CONVERTED INTO HYDROGEN SULPHIDE, SAID HYDROGEN-RICH GAS MIXTURE CONSTITUTING THE SOLE SOURCE OF HYDROGEN SUPPLIED TO SAHD HYDROFINING STAGE, SEPARATING HYDROGEN SULPHIDE FROM THE PRODUCTS OF SAID SECOND DESULPHURISATION ZONE, AND RECOVERING A SUBSTANTIALLY SULPHUR-FREE PRODUCTS FROM BOTH SAID DESULPHURISATION ZONES.
US213176A 1950-03-15 1951-02-28 Catalytic desulfurization of petroleum hydrocarbons Expired - Lifetime US2672433A (en)

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Cited By (7)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US2833697A (en) * 1953-10-23 1958-05-06 Basf Ag Desulfurization of crude oils by catalytic high-pressure hydrogenation
US2880166A (en) * 1955-10-19 1959-03-31 Phillips Petroleum Co Improving color and color stability of kerosene by combined autofining and hydrofining
US2883337A (en) * 1954-06-21 1959-04-21 Union Oil Co Process for hydrodesulfurizing hydrocarbons employing an impure hydrogen stream
US2901417A (en) * 1954-05-17 1959-08-25 Exxon Research Engineering Co Hydrodesulfurization of a coked hydrocarbon stream comprising gasoline constituents and gas oil constituents
US2951032A (en) * 1956-02-16 1960-08-30 Union Oil Co Hydrocarbon desulfurization process
US2983669A (en) * 1958-12-30 1961-05-09 Houdry Process Corp Hydrodesulfurization of selected gasoline fractions
US3250698A (en) * 1962-10-02 1966-05-10 British Petroleum Co Autofining of petroleum hydrocarbons

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* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
DE1025547B (en) * 1954-09-14 1958-03-06 Metallgesellschaft Ag Process for the catalytic refining of benzene or other hydrocarbons resulting from the coking, smoldering or gasification of hard coal
DE1140918B (en) * 1957-03-23 1962-12-13 Metallgesellschaft Ag Process for the production of the purest benzene and toluene by hydrogen refining and azeotropic distillation
US3429803A (en) * 1966-08-01 1969-02-25 Universal Oil Prod Co Method for converting hydrocarbons

Citations (4)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US2417308A (en) * 1943-04-12 1947-03-11 Union Oil Co Desulphurization and hydroforming
US2486361A (en) * 1944-10-20 1949-10-25 Union Oil Co Catalytic conversion of hydrocarbons
US2567252A (en) * 1949-07-20 1951-09-11 Anglo Iranian Oil Co Ltd Refining of hydrocarbons
US2573726A (en) * 1947-06-30 1951-11-06 Anglo Iranian Oil Co Ltd Catalytic desulphurisation of naphthas

Patent Citations (4)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US2417308A (en) * 1943-04-12 1947-03-11 Union Oil Co Desulphurization and hydroforming
US2486361A (en) * 1944-10-20 1949-10-25 Union Oil Co Catalytic conversion of hydrocarbons
US2573726A (en) * 1947-06-30 1951-11-06 Anglo Iranian Oil Co Ltd Catalytic desulphurisation of naphthas
US2567252A (en) * 1949-07-20 1951-09-11 Anglo Iranian Oil Co Ltd Refining of hydrocarbons

Cited By (7)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US2833697A (en) * 1953-10-23 1958-05-06 Basf Ag Desulfurization of crude oils by catalytic high-pressure hydrogenation
US2901417A (en) * 1954-05-17 1959-08-25 Exxon Research Engineering Co Hydrodesulfurization of a coked hydrocarbon stream comprising gasoline constituents and gas oil constituents
US2883337A (en) * 1954-06-21 1959-04-21 Union Oil Co Process for hydrodesulfurizing hydrocarbons employing an impure hydrogen stream
US2880166A (en) * 1955-10-19 1959-03-31 Phillips Petroleum Co Improving color and color stability of kerosene by combined autofining and hydrofining
US2951032A (en) * 1956-02-16 1960-08-30 Union Oil Co Hydrocarbon desulfurization process
US2983669A (en) * 1958-12-30 1961-05-09 Houdry Process Corp Hydrodesulfurization of selected gasoline fractions
US3250698A (en) * 1962-10-02 1966-05-10 British Petroleum Co Autofining of petroleum hydrocarbons

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