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
  • C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
  • C10G11/14—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts
  • C10G11/18—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils with preheated moving solid catalysts according to the "fluidised-bed" technique
• • 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
  • C10G11/00—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils
  • C10G11/02—Catalytic cracking, in the absence of hydrogen, of hydrocarbon oils characterised by the catalyst used
  • C10G11/04—Oxides
  • C10G11/05—Crystalline alumino-silicates, e.g. molecular sieves

Definitions

• the present invention relates to a process for fluid catalytic cracking (FCC) of hydrocarbon feeds in a downflow reactor using a specified cracking catalyst.
• FCC fluid catalytic cracking
• FCC processes are well known. In the more usual FCC processes employing riser reactors the catalyst and the hydrocarbon feed flow upward, while in FCC processes employing downflow reactors the catalyst and the hydrocarbon feed flow downward.
• downflow reactors do not display large differences in velocity and catalyst density between the centre and the wall of the reactor.
• the difference in velocity between the catalyst flow and the hydrocarbon flow in these reactors is smaller than in riser reactors.
• the slip factor of downflow reactors generally is about 1.
• the present invention provides a process for cracking hydrocarbon feeds which combines the advantages of downflow and riser reactors: minimal overcracking and high conversion of the higher-boiling fraction.
• the process comprises the following steps: a) atomizing and injecting a hydrocarbon feedstock into the top portion of a tubular downflow reactor and contacting this hydrocarbon feedstock with a catalyst having an AAI of at least 3.5, b) separating reaction products and spent catalyst at the bottom of said downflow reactor, c) treating the spent catalyst with steam, d) regenerating the spent catalyst in a regeneration zone, and e) recycling the regenated catalyst to the downflow reactor.
• the AAI is a measure of the accessibility of the catalyst pores to large, often high-molecular weight compounds and can be determined according to the method described in non pre-published European patent application No. 01202147.3 filed on June 5, 2001 , which application is incorporated by reference. This method involves adding the porous material to a stirred vessel containing large, preferably rigid, and often high-molecular weight compounds dissolved in a solvent and periodically analysing the concentration of these compounds in the solution. The relative concentration of the large compounds (in %) can be plotted against the square root of time (in minutes). The AAI is defined as the initial slope of this plot. The higher the AAI value, the more accessible the catalyst pores are.
• the AAI is not equivalent to the pore volume of a catalyst.
• the AAI deals with the accessibility of this pore volume, e.g. the size of the pore entrance.
• catalysts with a high pore volume can have low AAI values if the pore entrances are narrow.
• a hydrocarbon feedstock is atomized and injected into the top portion of a tubular downflow reactor, thereby contacting this hydrocarbon feedstock in the absence of added hydrogen with a hot, fluidised stream of catalyst having an AAI of at least 3.5.
• the spent catalyst having coke and hydrocarbonaceous material deposited thereon, is separated from the reaction products.
• the hydrocarbonaceous material is stripped from the spent catalyst by treatment with steam.
• the coke is removed from the spent catalyst during the regeneration step, involving combustion of the coke in an oxygen-containing atmosphere at a temperature of about 600- 850°C, preferably 650-750°C.
• the regenerated catalyst is recycled to the downflow reactor.
• the catalyst-oil contact time preferably is 0.5 to 5 seconds, more preferably 0.5 to 4 seconds, and even more preferably 1 to 3 seconds.
• the temperature at the reactor outlet preferably is between 450 and 700°C, more preferably between 500 and 600°C.
• the catalyst/oil ratio preferably is between 2 and 15.
• the spent catalyst is continuously removed from the reaction zone and made up with catalyst essentially free of coke resulting from the regeneration zone.
• fresh catalyst is regularly added to the process. If desired, part of the catalyst inventory can be withdrawn and replaced by fresh catalyst to adjust, e.g., the activity, selectivity or metal contamination of the circulating catalyst inventory.
• the fluidisation of the catalyst with various gas streams allows the transport of the catalyst between the reaction zone and the regeneration zone.
• the AAI of the catalyst to be used in the process according to the invention is at least 3.5, preferably at least 5.0, more preferably at least 6.0.
• the maximum AAI value depends on the required physical properties, such as apparent bulk density and friction strength.
• the catalyst preferably comprises 10-60 wt.% of a solid acid, 0-50 wt.% of alumina, 0-40 wt.% of silica, and the balance kaolin. More preferably, the catalyst comprises 20-50 wt.% of solid acid, 5-40 wt.% of alumina, 5-25 wt.% of silica, and the balance kaolin. Most preferably, the catalyst comprises 25-45 wt.% of solid acid, 10-30 wt.% of alumina, 5-20 wt.% of silica, and the balance kaolin.
• the catalyst may comprise solid acid, matrix, and/or any other component commonly used in FCC catalysts such as metal passivating agents.
• the matrix typically contains silica, alumina, silica-alumina, and/or clay.
• a preferred clay is kaolin.
• the solid acid can be a zeolite, e.g., a ZSM-type zeolite such as ZSM-5 or a faujasite-type zeolite, a silicoaluminophosphate (SAPO), an aluminophosphate (ALPO), or a combination thereof.
• the solid acid is a zeolite, more preferably a faujasite-type zeolite.
• the zeolite is optionally ultrastabilised and/or rare earth exchanged, e.g. zeolite Y, zeolite USY, zeolite REY, or zeolite REUSY.
• the rare earth content of the zeolite preferably is below 16 wt%.
• the micropore volume of the catalyst preferably is at least 0.050 ml/g, whereas the external surface area preferably is at least 100 m 2 /g.
• the first method comprises mixing the catalyst components or precursors thereof in an aqueous slurry to form a precursor mixture, adding a pore-forming agent to this mixture, followed by shaping, e.g. spray-drying.
• the pore-forming agent controls the porosity of the catalyst.
• a preferred pore- forming agent is a water-soluble carbohydrate, e.g., sucrose, maltose, cellobiose, lactose, glucose, or fructose. These pore-forming agents can be readily removed after the catalyst preparation. Thermogravimetric analyses indicate that the pore-forming agent can be removed to less than 5 wt.% remaining in the catalyst.
• the catalyst components or precursors thereof are mixed in an aqueous slurry to form a precursor mixture, the mixture is fed to a shaping apparatus and shaped to form particles, in which process just before being fed to the shaping apparatus the mixture is destabilised, i.e. its viscosity is increased.
• this method involves feeding suspended catalyst components or precursors thereof from one or more vessels (the "holding vessels") via a so-called pre-reactor to a shaping apparatus.
• the catalyst precursor mixture is destabilised.
• a destabilised mixture is defined as a mixture whose viscosity is higher after leaving the pre-reactor (and before shaping) than before entering the pre-reactor. The viscosity increase is due to induced polymerisation or gelling of catalyst binder material in the pre-reactor.
• the viscosity is typically increased from a level of about 1-100 Pa s at a shear rate of 0.1 s "1 before entering the pre-reactor to a level of about 50-1 ,000 Pa s or higher at a shear rate of 0.1 s "1 after leaving the pre-reactor. In any case, it is preferred to induce a viscosity increase of at least 10 Pa s, more preferably at least 50 Pa s, and most preferably at least 100 Pa-s (measured at a shear rate of 0.1 s "1 ).
• the viscosity is increased from a level of about 1-50 Pa.s at a shear rate of 0.1 s "1 before entering the pre-reactor to a level of about 50-500 Pa s at a shear rate of 0.1 s "1 after leaving the pre-reactor.
• the viscosity can be measured by standard rheometers, such as plate-and-plate rheometers, cone- and-plate rheometers or bop-and-cup rheometers.
• Destabilisation of the catalyst precursor mixture is performed in the pre-reactor just before feeding the mixture to the shaping apparatus.
• the time period involved i.e. the time which elapses between the start of the destabilisation and the shaping, depends on the exact configuration of the pre-reactor and on the time needed thereafter for the destabilised mixture to reach the shaping apparatus. Time periods of up to half an hour are possible, but may be less preferred for economical reasons. Preferred is a time period of less than 300 seconds. A more preferred time period is less than 180 seconds.
• Destabilisation can be performed by temperature increase, pH increase or pH decrease, and/or the addition of gel-inducing agents such as salts, phosphates, sulphates, and (partially) gelled silica.
• a suitable shaping method is spray-drying. For more details concerning this method we refer to non pre-published European patent application No. 01202146.5.
• This Example compares the performance of a conventional catalyst in a downflow and a riser reactor.
• a conventional equilibrium catalyst was evaluated in two distinct pilot units, one comprising a downflow reactor and the other comprising a conventional riser reactor. Both units operated at the same reaction temperature.
• the properties of the gas oil used are listed in Table 1.
• Table 2 displays the results of the cracking process using the riser and the downflow reactor at constant coke production. From these results it follows that the use of a downflow reactor leads to improved conversion levels and improved selectivity to C 3 olefins, as well as to improved hydrogen selectivity. However, the bottoms conversion in the downflow reactor is lower than in the riser reactor. Table 2
• a catalyst was prepared in the following way: A silica hydrosol was prepared by the controlled neutralisation under acidic pH of a sodium silicate solution by diluted sulfuric acid. To the freshly prepared hydrosol were added, sequentially and under thorough agitation, powdered kaolin, an acidic suspension of a boehmite-type alumina, and an acidic suspension of REY-zeolite. The resulting precursor suspension had a solids content of 20 wt%.
• the precursor mixture was subsequently fed to a spray-dryer and catalyst microspheres were recovered.
• the microspheres were re-suspended in ammoniated water and filtered under reduced pressure.
• the so-formed filter cake was twice exchanged with an ammonium sulfate solution at 45°C and washed three times with water at the same temperature.
• the catalyst particles were dried in an oven under circulating air at 110°C for 16 hours, which yielded the fresh sample EC2.
• EC2 was composed of 40 wt% of ultrastabilised Y-zeolite with a SAR of 5.5 and exchanged to reach 5 t% rare earth oxides (RE 2 O 3 ); 30 wt% silica-alumina matrix; and 30 wt% kaolin.
• the physical properties of this catalyst are displayed in Table 3.
• the catalyst of this Example was prepared using exactly the same procedure as that of Comparative Example 2, except that - as taught in Brazilian PI BR 9704925-5A - sucrose was added to the precursor mixture. This resulted in catalyst E1.
• the physical properties of this catalyst are displayed in Table 3.
• BET is the well-known BET surface area
• MiPV is the micropore volume
• MSA the mesopore (20-500A) surface area, all determined by N 2 adsorption (t-plot method).
• ABD stands for the Apparent Bulk Density, which is defined as the mass of catalyst per unit of volume in a non-compacted bed. The ABD is measured after filling a gauged cylinder of fixed, pre-determined volume without compaction of the bed.
• the AAI was determined by preparing a 1 I solution of 15 g Kuwait VGO in toluene by heating a Kuwait VGO feed to 70°C in an oven. 15 g of the warm Kuwait VGO were suspended in 200 ml warm toluene. The mixture was well stirred and adjusted to 1 litre with toluene. The solution was stored in the dark. 50.00 g of this solution were added to a 100 ml beaker (glass) connected to a peristaltic pump and a detector by way of tubes. The solution was stirred with a propeller stirrer at 400 rpm and the peristaltic pump was set at 21 g/min. A spectrophotometer was used as detector. This spectrophotometer was set to zero using a toluene solution.
• AAI Akzo Accessibility Index
• catalysts EC2 and E1 were hydrothermally deactivated using a 100% steam atmosphere at 788°C for 5 hours in order to simulate the equilibrium state.
• the resulting deactivated catalysts are called EC2D and E1 D, respectively.
• the deactivated samples were tested in the same downflow reactor-containing unit and under the same conditions as in Comparative Example 1. The results at the same conversion levels and the same coke levels are listed in Table 4.
• the unit inventory was 2 kg and the gas oil flow rate was 1.7 kg/h.
• the operating conditions were: reaction pressure 0.1 kgf/cm 2 g, contact time 2 seconds, temperature at the reactor exit 540°C and in the stripper 500°C.
• the catalyst/oil ratio (wt/wt) was varied in the range 6-9 by altering the feed temperature in the adiabatic reactor. Table 4
• Example 1 On using the catalyst of Example 1 it is possible to operate in a downward flow reactor without conversion loss at lower catalyst-to-oil ratios than those recommended in the prior art.

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• Chemical & Material Sciences (AREA)
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• Engineering & Computer Science (AREA)
• Chemical Kinetics & Catalysis (AREA)
• General Chemical & Material Sciences (AREA)
• Organic Chemistry (AREA)
• Crystallography & Structural Chemistry (AREA)
• Production Of Liquid Hydrocarbon Mixture For Refining Petroleum (AREA)
• Catalysts (AREA)

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• 2002
  • 2002-05-24 CA CA002449642A patent/CA2449642A1/en not_active Abandoned
  • 2002-05-24 CN CNB028115414A patent/CN1225522C/zh not_active Expired - Fee Related
  • 2002-05-24 JP JP2003503737A patent/JP2004535489A/ja active Pending
  • 2002-05-24 WO PCT/EP2002/005745 patent/WO2002100983A2/en active Application Filing
  • 2002-05-24 KR KR10-2003-7015950A patent/KR20040012899A/ko not_active Application Discontinuation
  • 2002-05-24 EP EP02743095A patent/EP1392796A2/en not_active Withdrawn
  • 2002-05-24 US US10/480,966 patent/US20040206669A1/en not_active Abandoned
  • 2002-05-24 BR BR0210168-8A patent/BR0210168A/pt not_active Application Discontinuation

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