Title of Invention

A PROCESS FOR THE FLUID CATALYTIC CRACKING OF HYDROCARBONS

Abstract

ABSTRACT A PROCESS FOR THE FLUID CATALYTIC CRACKING OF HYDROCARBONS 1949/CHENP/2003 The present invention relates to a process for the fluid catalytic cracking of hydrocarbons comprising 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 Akzo Accessibility Index (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 recycling the regenerated catalyst to the downflow reactor; wherein the catalyst comprises 10- 60 wt.% of a solid acid, 0-50% wt.% of alumina, 0-40% wt.%) of silica, and the balance kaolin.

Full Text

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 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. In riser reactors solids flow upward due to the lift caused by the ascending vaporised feed. However, the velocity of the hydrocarbon feed is lower near the wall than it is near the centre of the reactor. Therefore, the catalyst will move more slowly near the reactor wall than near the centre, resulting in a slower moving area with a high catalyst density near the wall and a low-resistance path of ascending feed near the centre. Hence, the feed mainly flows through the centre, whereas the catalyst is mainly located near the walls. The resulting flow pattern is called core-annulus. Furthermore, the upward flow of solid catalyst and hydrocarbon vapour in riser reactors opposes gravity, resulting in a catalyst flow that is significantly slower than the much lighter hydrocarbon flow. The ratio of feed velocity to catalyst velocity, i.e. the slip factor, generally is about 2-3. This results in backmixing of the catalyst, leading to longer residence times of the catalyst and the occurrence of undesirable secondary reactions (overcracking). In contrast to riser reactors, downflow reactors do not display large differences in velocity and catalyst density between the centre and the wall of the reactor. Furthermore, as the catalyst particles do not oppose gravity, 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. Consequently, backmixing is largely avoided, the catalyst is distributed more evenly across the entire reactor, and the effective contact time of the catalyst and the feed in a downflow reactor is less than in a riser reactor. Although this reduces the formation of by-products, it also results in a decrease In the conversion of mainly the larger, higher-boiling compounds. For prior art publications dealing with cracking units comprising downflow reactors, reference is made to US 5.449,496, US 5,582,712, US 6,099,720, US 5,660,716, US 5,951,850, and EP 0 254 333. Of these publications only a few focus on optimising the process by way of using a specified catalyst. In U.S. 5,660,716 use is made of a low acidity catalyst. It is recommended to use it in conjunction with high temperatures and high catalyst-to-oil ratios to obtain acceptable conversion levels. A similar teaching - including the use of catalyst-to-oil ratios of 25 to 80 w/w % - is contained in U.S. 5,951,850, in which it is recommended to use a catalyst containing a zeolite having a unit cell size of at most 24.50 Angstroms. The high catalyst-to-oil ratio may jeopardize the performance of the unit as far as catalyst separation, stripping, and regeneration capabilities are concerned. Moreover, wear of the equipment caused by the catalyst may become critical. In sum, the teachings of the prior art tend into the direction of using low activity catalysts in conjunction with high temperatures and high catalyst-to-oil ratios to compensate for the lower catalyst activity. 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 use of a downflow reactor minimises overcracking, while the high accessibility of the cracking catalyst facilitates high conversions, even for high- boiling fractions in the hydrocarbon feed. The AAI is a measure of the accessibility of the catalyst pores to large, often high-molecular weight compounds and can be detemriined 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. If the pores of the cracking catalyst are highly accessible to even the higher- boiling fractions of the hydrocarbon feed, the feed molecules will diffuse quickly through the pores and optimum use is made of the active sites present in the catalyst pores. Hence, high conversions can be reached with such catalysts. It is emphasised that the Ml 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. Hence, catalysts with a high pore volume can have low AAI values if the pore entrances are narrow. According to the process according to the invention, 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. Next, 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- SSCC, preferably BSO-TSCC. Finally, 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. To make up for catalyst losses, 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 catalyst 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. Preferably, 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^/g. Suitable methods for the preparation of such highly accessible catalysts include the methods disclosed in Brazilian patent publication BR PI 9704925-5A and in non pre-published European patent application No. 01202146.5, filed on June 5, 2001, which applications are both incorporated by reference. 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. According to the second method, 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. VIore in particular, this method involves feeding suspended catalyst components or precursors thereof from one or more vessels (the "holding /essels") via a so-called pre-reactor to a shaping apparatus. In this pre-reactor he catalyst precursor mixture is destabilised. n this specification a destabilised mixture is defined as a mixture whose fiscosity is higher after leaving the pre-reactor (and before shaping) than before jntering the pre-reactor. The viscosity increase is due to induced polymerisation )r gelling of catalyst binder material in the pre-reactor. The viscosity is typically ncreased from a level of about 1-100 Pas at a shear rate of 0.1 s'^ before sntering the pre-reactor to a level of about 50-1,000 Pas or higher at a shear ate of 0.1 s"^ after leaving the pre-reactor. In any case, it is preferred to induce I viscosity increase of at least 10 Pa-s, more preferably at least 50 Pas, and lost preferably at least 100 Pas (measured at a shear rate of 0.1 s'^). Preferably, the viscosity is increased from a level of about 1-50 Pa.s at a shear ate of 0.1 s'^ before entering the pre-reactor to a level of about 50-500 Pas at a shear rate of 0.1 s'^ 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. EXAMPLES Comparative Example 1 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 1 Physical and chemical properties of the gas oil used "API i^^e Density 20/4°C (g/ml) 0.9386 Viscosity (ASTM D445) (cSt) 268 Aniline point f C) 80.8 Basic Nitrogen (ppm) 961 Concarbon Residue (wt%) 0.38 Initial Boiling Point, IBP (°C) 309 Final Boiling Point, FBP (°C) 602 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 C3 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 Riser reactor 550 Downflow reactor Reaction Temperature (°C) 550 Conversion {wt%) 72.6 74.8 Catalyst/Oil Ratio, CTO (wt/wt) 7.8 8.7 Delta Coke (Coke/CTO, wt%) 1.13 1.02 Coke (wt%) 8.8 8.8 Fuel Gas (wt%) 4.8 4.8 Hydrogen (wt%) 0.60 0.17 LPG {wt%) 17.9 20.4 Propene {wt%) 4.84 6.63 Gasoline (wt%) 41.0 40.8 LCO (wt%) 15.9 12.3 Bottoms (wt%) 11.6 13.0 Comparative Example 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. Finally, 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 wt% rare earth oxides (RE2O3); 30 wt% silica-alumina matrix; and 30 wt% kaolin. The physical properties of this catalyst are displayed in Table 3. Example 1 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. Table 3 BET (m'/g) MiPV (ml/g) MSA (m'/g) ABD (kg/dm') AAI EC2 E1 287 362 0.103 0.115 66 110 0.71 0.70 2.0 6.0 In this Table, BET is the well-known BET surface area, MiPV is the micropore volume, and MSA the mesopore (20-500A) surface area, all determined by Na 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. Next, 1 g of a 53-75 microns sieve fraction of the catalyst was added to the Kuwait VGO in toluene solution. Once per second the asphaltene concentration was measured by spectrophotometry at a wavelength of 560 nm. After 5 minutes, the measurement was stopped and the relative absorbance was plotted versus the square root of time. The slope, i.e. the Akzo Accessibility Index (AAI), was determined. Example 2 Portions of catalysts EC2 and El 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 E1D, 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^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 EC2D E1D Equal Conversion (wt%) 77.0 77.0 Catalyst /Oil Ratio (wt/wt) 8.6 6.0 Coke (wt%) 8.9 7.8 Fuel Gas (wt%) 4.0 3.2 Hydrogen (wt%) 0.10 0.04 LPG (wt%) 19.0 17.4 Propene (wt%) 5.1 4.5 Gasoline (wt%) 45.1 48.6 LCD (wt%) 12.2 12.6 Bottoms (wt%) 10.8 10.4 Equal Coke (wt%) 8.0 8.0 Catalyst /Oil Ratio (wt/wt) 6.4 6.9 Conversion (vvt%) 72.2 79.7 Fuel Gas (wt%) 3.6 3.3 Hydrogen(wt%) 0.10 0.04 LPG (wt%) 17.0 19.3 Propene (wt%) 4.4 4.9 Gasoline (wt%) 43.6 49.0 LCO (wt%) 13.2 12.0 Bottoms (wt%) 14.7 8.4 From this Table it is clear that a process using a combination of a downflow reactor and a catalyst with an AAI of at least 3.5 results in high conversion levels and gasoline yields, combined with high bottoms conversion. These results were obtained using reaction temperatures and catalyst-to-oil ratios which are normally employed in industrial riser reactors. Moreover, the feed used was one having a high basic nitrogen content. 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. The results also show a tendency towards improvements in coke, fuel gas and gasoline selectivity. At constant coke the synergism between the use of a downflow reactor and a catalyst having an AAI of at least 3.5 is especially beneficial. Fuel gas and hydrogen yields are reduced and light olefins are increased. Compared to the base case bottoms conversion is increased too. This shows that the disadvantage in respect of bottoms conversion for downflow operations observed in the base case - see Table 2 - may be fully compensated by the rocess according to the invention. inally, it has been observed that the strippability of catalysts having an AAI of t least 3.5 is greatly improved in comparison with prior art catalysts not having uch high accessibility. We claim: 1. A process for the fluid catalytic cracking of hydrocarbons comprising 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 Akzo Accessibility Index (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 regenerated catalyst to the downflow reactor; wherein the catalyst comprises 10-60 wt.% of a solid acid, 0-50% wt.% of alumina, 0-40% wt.% of silica, and the balance kaolin. 2. The process as claimed claim 1 wherein the catalyst has an AAI of at least 5.0. 3. The process as claimed in claim 2 wherein the catalyst has an AAI of at least 6.0. 4. The process as claimed in claim 1 wherein the catalyst has been obtained by mixing the catalyst components or precursors thereof in an aqueous mixture, adding a pore-forming agent to this mixture, followed by shaping. 5. The process as claimed in claim 4 wherein the pore-forming agent is a water- soluble carbohydrate. 6. The process as claimed in claim 1 wherein the catalyst comprises 20-50 wt.% of solid acid, 5-40 wt.% of alumina, 5-25 wt.% of silica, and the balance kaolin. 7. The process as claimed in claim 1 wherein the catalyst comprises 25-45 wt.% of solid acid, 10-30 wt.%» of alumina, 5-20 wt.% of silica, and the balance kaolin. 8. The process as claimed in claim 6 wherein the solid acid is selected from the group consisting of ZSM-type zeolites, faujasite-type zeolites, siiicoaluminophosphate (SAPO), aluminophosphate (ALPO), and combinations thereof 9. The process as claimed in claim 8 wherein the solid acid is a rare earth exchanged zeolite. 10. The process as claimed in claim 9 wherein the rare earth content of the zeolite is below 16 wt%. 11. The process as claimed in claim 1 wherein the catalyst has been obtained by combining catalyst components or precursors thereof in an aqueous medium to form a catalyst precursor mixture, feeding the mixture to a shaping apparatus, and shaping the mixture to form particles, in which process just before being fed to the shaping apparatus the mixture is destabilized.

Documents:

1949-chenp-2003 abstract.pdf

1949-chenp-2003 assignment.pdf

1949-chenp-2003 claims duplicate.pdf

1949-chenp-2003 claims.pdf

1949-chenp-2003 correspondence - others.pdf

1949-chenp-2003 correspondence - po.pdf

1949-chenp-2003 description (compelet) duplicate.pdf

1949-chenp-2003 description (compelet).pdf

1949-chenp-2003 form - 1.pdf

1949-chenp-2003 form - 18.pdf

1949-chenp-2003 form - 26.pdf

1949-chenp-2003 form - 3.pdf

1949-chenp-2003 form - 5.pdf

1949-chenp-2003 form - 6.pdf

1949-chenp-2003 pct search report.pdf

1949-chenp-2003 pct.pdf

1949-chenp-2003 petition.pdf