Classifications

• • C—CHEMISTRY; METALLURGY
  • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
  • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
  • C10G65/00—Treatment of hydrocarbon oils by two or more hydrotreatment processes only
  • C10G65/02—Treatment of hydrocarbon oils by two or more hydrotreatment processes only plural serial stages only
  • C10G65/12—Treatment of hydrocarbon oils by two or more hydrotreatment processes only plural serial stages only including cracking steps and other hydrotreatment steps
• • 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
  • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
  • C10G2400/04—Diesel oil

Definitions

• the present invention relates to the field of fuels for internal combustion engines. It relates more particularly to the manufacture of a fuel for a compression ignition engine and the fuel thus obtained.
• diesel cuts whether they come from the direct distillation of crude oil or whether they come from a conversion process such as catalytic cracking, still contain significant amounts of aromatic compounds, nitrogen compounds and sulfur compounds.
• the fuel usable in engines must contain a quantity of sulfur lower than 500 parts per million in weight (ppm). In the vast majority of these countries there are currently no standards imposing a maximum content of aromatic compounds and nitrogen.
• diesel cuts come either from direct distillation of crude or from cracking catalytic: i.e. cuts of light distillates (Anglo-Saxon initials LCO for Light Cycle Oil), heavy fraction cuts (Anglo-Saxon initials HCO for Heavy Cycle Oil), or another conversion process (coking, visbreaking, residue hydroconversion etc.) or even diesel from distillation of aromatic or naphthenoaromatic crude oil such as Hamaca, Zuata, EI Pao. It is particularly important to produce an effluent directly and fully recoverable as a very high quality fuel cutter.
• the present invention differs from the prior art in that it combines hydrocracking with hydrogenation.
• Such a combination has already been described for the treatment of heavy loads, for example in patent FR-A-2,600,669.
• the charge treated contains at least 50% by weight of constituents boiling above 375 ° C. and the aim of the process is to convert at least 70% vol. of these heavy constituents into constituents with boiling points below 375 ° C.
• This two-step process essentially involves significant hydrogenation or controlled aromatic compounds - depending on the content of aromatic compounds that we want to obtain in the final product, then a hydrocracking intended to open the naphthenes produced in the first stage so as to form paraffins.
• These charges are treated with hydrogen in the presence of catalysts, this treatment allows to hydrogenate the aromatic compounds present in the feed, it also allows to carry out hydrodesulphurization and hydrodenitrogenation simultaneously.
• the operating conditions of the hydrogenation (or hydrotreatment) are the following: the space speed (V.V.H.) is between 0.1 and 30 volumes of liquid feed per volume of catalyst and per hour and preferably between 0.2 and 10; the inlet temperature to the reactor is between 250 and 450 ° C and preferably between 320 and 400 ° C; pressure at the reactor is between 0.5 and 20 MPa and preferably between 4 and 15 MPa; recycling of pure hydrogen is between 100 and 2,500 Nm3 / m3 of charge and preferably between 200 and 2100 Nm3 / m3, and even more advantageously less than 2000 Nm3 / m3.
• the hydrogen consumption in the process can go up to about 5% by weight of the load (0.5-4.5% in general).
• the hydrogenation catalyst comprises, on an amorphous mineral support, at least one metal or compound of metal from group VIB of the periodic table of elements such as molybdenum or tungsten, in an amount expressed by weight of metal relative to the weight of the finished catalyst of between 0.5 and 40% and preferably between 2 to 30%, at least one metal or compound of non-noble metal from group VIII of said periodic classification such as nickel, cobalt or iron in an amount expressed by weight of metal relative to the weight of the finished catalyst of between 0.01 and 30% and preferably between 0.1 and 10%, of phosphorus or at least one phosphorus compound in an amount expressed by weight of pentoxide of phosphorus relative to the weight of the support between 0.001 and 20%.
• group VIB of the periodic table of elements such as molybdenum or tungsten
• the catalyst can also contain boron or at least one boron compound in an amount expressed by weight of boron trioxide relative to the weight of the support of between 0.001 and 10%.
• the amorphous mineral support will be, for example, alumina or silica-alumina. According to a particular form of the invention, cubic gamma alumina will be used which preferably has a specific surface of approximately 50 to 500 m 2 / g.
• the hydrogenation catalyst used in the present invention is preferably subjected to a sulfurization treatment making it possible to transform, at least in part, the metallic sulphide species before their contact with the load to be treated.
• This activation treatment by sulfurization is well known to those skilled in the art and can be carried out by any method already described in the literature.
• a conventional sulfurization method well known to those skilled in the art consists in heating the catalyst in the presence of hydrogen sulfide or a hydrogen precursor sulfurized at a temperature between 150 and 800 ° C, preferably between 250 and 600 ° C, generally in a crossed-bed reaction zone.
• hydrogen sulfide precursor within the meaning of the present description means any compound capable of reacting, under the operating conditions of the reaction for give hydrogen sulfide.
• the hydrogenated products coming from the first stage may or may not undergo a treatment chosen from the group formed by gas-liquid separations and distillations.
• the liquid phase then undergoes hydrocracking according to step b) of the present invention.
• the operating conditions for hydrocracking are as follows: the space velocity (V.V.H.) is approximately 0.1 to 30 charge volumes liquid per volume of catalyst and per hour and preferably between 0.2 and 10, the inlet temperature to the reactor is between 250 to 450 ° C and preferably between 300 and 400 ° C; the pressure in the reactor is between 0.5 and 20 MPa and preferably between 4 and 15 MPa and even more preferably between 7 and 15 MPa; the recycling of pure hydrogen is between 100 to 2200 Nm3 / m3 of charge. Under these conditions, the conversion is adjusted according to the cetane number. and other properties (density, T95 ...) to obtain. Total conversion (hydrocracking b) + that obtained during the hydrogenation step a)) may be greater than 50% or less than 50% (5-50% for example) depending on the cut to be treated.
• the catalyst of the second stage generally comprises at least one zeolite, at least one support and at least one hydro-dehydrogenating function.
• An acidic zeolite is particularly advantageous in this type of embodiment, for example a zeolite of faujasite type, and preferably a Y zeolite, will be used.
• the zeolite content by weight is between 0.5 and 80% and preferably between 3 and 50 % relative to the finished catalyst.
• a zeolite Y with a crystalline parameter 24.14 x 10 -10 m to 24.55 x 10 -10 m will be used.
• the catalyst contains at least one metal oxide or sulfide of group VIB such as molybdenum or tungsten in an amount expressed by weight of metal by relative to the weight of the finished catalyst of between 0.5 to 40% and at least one non-noble metal or compound of non-noble metal from group VIII such as nickel, cobalt or iron in an amount expressed by weight of metal relative by weight of the finished catalyst of between 0.01 and 20% and preferably between 0.1 and 10%.
• group VIB such as molybdenum or tungsten
• group VIII such as nickel, cobalt or iron
• the hydrocracking catalyst used in the present invention is preferably subjected to a sulfurization treatment making it possible to transform, at least in part, the metal species into sulphides before they are brought into contact with the feed to be treated.
• This sulfurization activation treatment is well known to those skilled in the art and can be carried out by any method already described in the literature.
• a conventional sulfurization method well known to those skilled in the art consists in heating the catalyst in the presence of hydrogen sulfide or of a hydrogen sulfide precursor at a temperature between 150 and 800 ° C., preferably between 250 and 600 ° C, generally in a crossed bed reaction zone.
• a particularly advantageous acidic zeolite HY is characterized by different specifications: a SiO 2 / Al 2 O 3 molar ratio of between 8 and 70 and preferably between 12 and 40: a sodium content of less than 0 , 15% by weight determined on the zeolite calcined at 1100 ° C; a crystalline parameter "a" of the elementary mesh comprised between 24.55 x 10 -10 m and 24.24 x 10 -10 m and preferably between 24.38 x 10 -10 m and 24.26 x 10 -10 m; a CNa capacity for taking up sodium ions, expressed in grams of Na per 100 grams of modified zeolite, neutralized then calcined, greater than 0.85; a specific surface area determined by the BET method greater than approximately 400 m 2 / g and preferably greater than 550 m 2 / g, a water vapor adsorption capacity at 25 ° C.
• a porous distribution comprising between 1 and 20% and preferably between 3 and 15% of the pore volume contained in pores with a diameter between 20 x 10 -10 m and 80 x 10 -10 m, the rest of the pore volume being mainly contained in the pores of diameter less than 20 x 10 -10 m.
• the Y-Na zeolite from which the HY zeolite is prepared has a SiO 2 / Al 2 O 3 molar ratio of between approximately 4 and 6; it should first be lowered the sodium content (weight) to a value of the order of 1 to 3% and preferably to less than 2.5%; the Y-Na zeolite also generally has a specific surface of between 750 m 2 / g and about 950 m 2 / g
• variants of preparations exist which generally follow the hydrothermal treatment of the zeolite with an acid treatment.
• the effluent obtained at the end of the hydrocracking is obviously fractionated to separate the light products (cracked), i.e. products boiling below 150 ° C in general, or even below 180 ° C or other temperature chosen by the refiner.
• the light products i.e. products boiling below 150 ° C in general, or even below 180 ° C or other temperature chosen by the refiner.
• the present invention makes it possible to obtain diesel cuts whose cetane number, and possibly the content of aromatic compounds, are improved in such a way that these cuts will be able to meet current and future specifications. These diesel cuts are directly marketable.
• the present invention makes the most of all the products contained in the cut of treated oil.
• the yield of recoverable products is close to 99% in relation to the quantity of hydrocarbons; unlike other conventional processes, there is no liquid or solid waste to incinerate.
• the diesel feedstocks to be treated are, for example, direct distillation diesel, fluid catalytic cracking gas oils (English initials FCC for Fluid Catalytic Cracking) or (LCO). They generally have an initial boiling point of at least 180 ° C and a final boiling point of at most 370 ° C.
• the weight composition of these expenses by family of hydrocarbons is variable according to the intervals.
• the paraffin contents are between 5.0 and 30.0%, in naphthenes between 5.0 and 40.0% and in compounds aromatic between 40.0 and 80.0%.
• Less aromatic fillers can be also treated with less than 40% aromatics and generally 20% to less than 40% aromatics, the naphthene contents being able to go up to 60%.
• the catalyst used in the hydrogenation stage has the following characteristics: nickel content in the form of oxides of 3%, a molybdenum content in the form of oxides of 16.5% and 6% phosphorus pentoxide on alumina.
• a catalyst is advantageously used, the support of which is alumina. This catalyst contains by weight 12% of molybdenum, 4% of nickel in the form of oxides and 10% of zeolite Y, this catalyst is described in example 2 of US Pat. No. 5,525,209. These catalysts are sulfurized by an n-hexane / DMDS + aniline mixture up to 320 ° C. After 3000 hours of continuous operation, no deactivation of the catalysts as described in the example was observed.
• the charge is treated in a pilot unit comprising two reactors in series, under the following conditions: the space speed in the two reactors is 0.29 volume of liquid charge per volume of catalyst and per hour, the temperature of entry into the first reactor is 380 ° C for hydrogenation and it is 390 ° C for hydrocracking, the pressure in the two reactors is 14 MPa. In each reactor, the hydrogen recycling is 2000 Nm 3 per m 3 of feed. The characteristics of the feed and of the 190 ° C. + product obtained after each step are recorded in Table 1, after the hydrocracking step and after distillation.
• the charge was treated in a pilot unit comprising two reactors in series, under the following conditions, the space speed in the two reactors is 0.25 volume of liquid charge per volume of catalyst and per hour, the entry temperature into the first reactor is 385 ° C for hydrogenation and in the second reactor, it is 375 ° C for hydrocracking, the pressure in the two reactors is 14 MPa. In each reactor, the hydrogen recycling is 2000 Nm 3 per m 3 of feed. The characteristics of the charges and of the products obtained after each step are recorded in Table 2.
• the feed was treated in a pilot unit comprising the two reactors in series in Example 1, under the following conditions, the space speed in the two reactors is 0.25 volume of liquid feed per volume of catalyst and per hour, the inlet temperature in the first reactor is 360 ° C. for the hydrogenation and in the second reactor, it is 367 ° C. for the hydrocracking, the pressure in the two reactors is 14 MPa. In each reactor, the hydrogen recycling is 2000 Nm 3 per m 3 of feed. The characteristics of the fillers and of the products obtained after each step are recorded in Table 3.
• the invention has two major advantages: it saves hydrogen since less hydrogenation is carried out to obtain the same index of cetane; it also allows the constitution of a reserve of aromatic compounds that it is also possible, if necessary, to hydrogenate in a subsequent hydrogenation stage, this which results in a potential for increasing the cetane number. This last case relates more particularly to diesel starting cuts with aromatic contents high (40-80% wt).
• the hydrogenation stage is carried out with any catalyst known hydrogenation, and in particular those containing at least one noble metal deposited on an amorphous support of refractory oxide (alumina for example).
• a catalyst preferred contains at least one noble metal (preferred platinum), at least one halogen (and preferably 2 halogens: chlorine and fluorine) and a matrix (preferred alumina)
• the step hydrogenation can be carried out on the total effluent leaving the hydrocracking stage, separation of the compounds 150- (or preferably 180-) then taking place after this hydrogenation.
• the hydrogenation step can also be carried out on the 150+ section (or 180+ depending on the fractionation chosen), possibly followed by a separation of compounds 150- (or 180-).
• the limit imposed by conventional advanced hydrogenation processes is set by the aromatic content. Once these aromatic compounds are all hydrogenated, we can no longer hope to increase the cetane number, on the other hand associating hydrocracking with hydrogenation, we can further increase the cetane by increasing the paraffin content of the cut.
• the combination according to the invention of hydrogenation and then hydrocracking stages makes it possible to obtain an index of high cetane, which could not have been obtained by the advanced hydrogenation used in the prior art.
• the sequence of processes that we propose here allows to exceed the limit imposed by advanced hydrogenation processes and to increase cetane number beyond any specification.

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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)
• Turbine Rotor Nozzle Sealing (AREA)

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• 1998
  • 1998-04-09 FR FR9804605A patent/FR2777290B1/fr not_active Expired - Lifetime
• 1999
  • 1999-04-09 KR KR1020007011223A patent/KR100601822B1/ko not_active IP Right Cessation
  • 1999-04-09 JP JP2000543542A patent/JP2002511516A/ja active Pending
  • 1999-04-09 DE DE69913673T patent/DE69913673T2/de not_active Revoked
  • 1999-04-09 EP EP99913363A patent/EP1070108B9/fr not_active Revoked
  • 1999-04-09 BR BR9909546-7A patent/BR9909546A/pt not_active Application Discontinuation
  • 1999-04-09 US US09/445,573 patent/US6814856B1/en not_active Expired - Fee Related
  • 1999-04-09 ES ES99913363T patent/ES2213358T3/es not_active Expired - Lifetime
  • 1999-04-09 WO PCT/FR1999/000817 patent/WO1999052993A1/fr not_active Application Discontinuation
• 2004
  • 2004-08-12 US US10/916,537 patent/US20050029161A1/en not_active Abandoned

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