WO2011025801A1 - Réduction des dibenzothiophènes à empêchement stérique dans un distillat de craquage catalytique en lit fluidisé issu d'une unité de craquage catalytique en lit fluidisé à double zone réactionnelle - Google Patents

Réduction des dibenzothiophènes à empêchement stérique dans un distillat de craquage catalytique en lit fluidisé issu d'une unité de craquage catalytique en lit fluidisé à double zone réactionnelle Download PDF

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Publication number
WO2011025801A1
WO2011025801A1 PCT/US2010/046569 US2010046569W WO2011025801A1 WO 2011025801 A1 WO2011025801 A1 WO 2011025801A1 US 2010046569 W US2010046569 W US 2010046569W WO 2011025801 A1 WO2011025801 A1 WO 2011025801A1
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catalyst
fraction
cracking
riser
fcc
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PCT/US2010/046569
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English (en)
Inventor
Stacey E. Siporin
George A. Swan, Iii
Bruce R. Cook
Steven S. Lowenthal
Michael A. Hayes
Michael W. Bedell
Steve Colgrove
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Exxonmobil Research And Engineering Company
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Publication of WO2011025801A1 publication Critical patent/WO2011025801A1/fr

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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
    • 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
    • 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
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/4081—Recycling aspects

Definitions

  • This invention relates to a process for producing low sulfur distillates such as low sulfur road diesel fuel.
  • Sulfur is found in refinery streams in a number of different forms including aliphatic and aromatic sulfur compounds; in the lower boiling naphtha streams, mercaptans, sulfides and thiophenes predominate and these can be removed easily by extractive or oxidative/extractive processes such as the commercially available MeroxTM process.
  • the sulfur compounds concentrated in the higher boiling distillate fractions is mainly in the form of aromatic heterocyclic compounds such as the thiophenes, benzothiophenes and
  • DBTs dibenzothiophenes
  • benzothiophenes At higher desulfurization severities, the more refractory sulfur compounds can be removed although with increased cost and with greater difficulty. Certain sulfur compounds are more difficult to remove than others.
  • the most difficult compounds to remove by hydroprocessing are the dibenzothiophenes and, of these, the substituted dibenzothiophenes tend to be less amenable to hydrodesulrurization than dibenzothiophene itself; this effect varies according to the extent and type of substitution in the dibenzothiophenes with the sterically-hindered alkyl dibenzothiophenes such as the 4,6-dialkyl dibenzothiophenes being the most refractory. See Chemistry of Catalytic
  • LCO light cycle oil
  • FCC fluid catalytic cracking
  • hydrodesulfurization catalysts Another costly option is hydrotreating the hydrocarbon feedstream to the FCC, which reduces the sulfur content but also alters the composition of the sulfur free hydrocarbons, especially of the high octane olefins which enter the gasoline fraction. This last option is also very costly due to the large (i.e., non-selective) volume of hydrocarbons required to be hydrotreated.
  • the commercial success of these additives has, however, been limited. Additionally, as most refineries need additional capital hardware in order to treat any additional SO x loadings in an FCC unit, this option can be very costly in most instances.
  • the present invention we propose a method for reducing the level of hindered alky 1-DBTs in the FCC middle distillate product by subjecting at least a portion of the heavy fraction of the LCO cracking product to a transalkylation regime in a fluid catalytic unit.
  • the process of the present invention may be carried out in an FCC unit which has at least two reaction zones. These two reaction zones may be incorporated into a single riser or reactor.
  • the transalkylation process of the present invention is preferably carried out in a secondary riser zone of a dual-riser FCC unit with the secondary riser fed only with the heavy LCO fraction containing the hindered DBTs.
  • This secondary riser can be fed with fresh or freshly regenerated catalyst circulating in the FCC catalyst inventory in order to make use of the high activity of the catalyst in this form.
  • dibenzothiophenes in the LCO fraction from an FCC process is reduced by passing the portion of the light cycle oil cracking product that contains the alkylated dibenzothiophenes, typically at least a portion of the LCO fraction boiling substantially in the range from about 500 to about 75O 0 F (260 to 400°C) ⁇ and more preferably, the fraction boiling substantially in the range from about 520 to about 75O 0 F (271 to 400 0 C) 5 is recycled to transalkylate or dealkylate the sterically-hindered alkyl DBTs in the FCC feed in the presence of a circulating fluid catalytic cracking catalyst.
  • the recycled portion of the LCO fraction boils substantially in the range from about 520 to about 68O 0 F (271 to 36O 0 C).
  • the reaction transferring the alkyl groups from the alkylated DBTs to the other species present is favored by temperatures which are lower relative to the cracking temperatures encountered in the cracking cycle.
  • the use of the secondary reaction zone or secondary riser is advantageous in that the conditions, for example, catalyst temperature, catalyst:oil ratio, and riser residence time, can be separately controlled to values appropriate for the desired transalkylation reactions without affecting the conditions prevailing in the main, or first, reaction zone of the FCC unit.
  • FIGURE 1 is a simplified schematic of a dual riser FCC unit in which the secondary riser is used to transalkylate a heavy fraction of the LCO cracking product.
  • FIGURE 2 is a graph showing the sulfur speciation of a typical light cycle oil.
  • FIGURE 3 is a graph showing the sulfur speciation of cracking products of a vacuum gas oil (“VGO”) with dibenzothiophene added in three concentrations.
  • VGO vacuum gas oil
  • FIG. 1 shows, in simplified form, by way of example, a preferred embodiment of dual-riser FCC unit which is suitable for carrying out the present process.
  • the FCC unit comprises a primary cracking riser (10) to which fresh cracking feed is introduced through line (1 1).
  • Riser (10) terminates inside reactor or disengager vessel (12) and terminates conventionally with a separation device to ensure rapid separation of the catalyst from the vaporous cracking products.
  • Separation systems such as the closed positive pressure cyclone systems as described, for example, in U.S. Patents Nos. 5,055,177; 5,039,397; 4,909,993; 4,654,060; 4,581,205; 4,502,947; negative pressure systems as disclosed in U.S. Patent No.
  • the vaporous cracking products pass from reactor vessel (12) to FCC main fractionation column (13) with its associated side columns (not shown) to separate the cracking products into light gases, LPG, and liquid products such as the naphtha fractions including a light naphtha and heavy naphtha, light cycle oil and a heavy fuel oil or slurry oil fraction, according to refinery specification.
  • the spent catalyst from stripper (14) at the bottom of the reactor passes through standpipe (15) fitted with slide valve (16) for controlling the flow of catalyst to regenerator (17).
  • the catalyst undergoes oxidative regeneration in the regenerator and then passes out of the regenerator vessel by way of standpipe (18) and slide valve (19). Downstream of slide valve (19), the standpipe divides into two branches which lead, respectively, to primary riser (10) by way of standpipe (20) and by way of standpipe (21) to secondary riser (22).
  • a portion of the recycled LCO fraction from the product fractionation section is brought by way of line (25) to feed pipe (26) which leads into the mix zone of a smaller, secondary riser (22).
  • the secondary riser (22) leads into the reactor vessel for recovery of the catalyst from both riser and for common recovery of the vapor cracking/reaction products.
  • FCC fluid catalytic cracking
  • LCO fluid catalytic cracking
  • conventional FCC catalysts may be used, for example, zeolite based catalysts with a faujasite cracking component as described in the seminal review by Venuto and Habib, Fluid Catalytic Cracking with Zeolite Catalysts, Marcel Dekker, New York 1979, ISBN 0-8247-6870-1 as well as in numerous other sources such as Sadeghbeigi, Fluid Catalytic Cracking Handbook, Gulf Publ. Co. Houston, 1995, ISBN 0-88415-290-1.
  • the fluid catalytic cracking process in which the heavy hydrocarbon feed containing the organosulfur compounds will be cracked to lighter products takes place by contact of a hydrocarbon-containing feed (also referred to herein as “heavy hydrocarbon feed”, “hydrocarbon feed”, or simply "feed) in a cyclic catalyst recirculation cracking process with a circulating fluidizable catalytic cracking catalyst inventory consisting of particles having a size ranging from about 20 to about 100 microns.
  • a hydrocarbon-containing feed also referred to herein as “heavy hydrocarbon feed”, “hydrocarbon feed”, or simply "feed
  • a circulating fluidizable catalytic cracking catalyst inventory consisting of particles having a size ranging from about 20 to about 100 microns.
  • the hydrocarbon feed is catalytically cracked in a first catalytic cracking zone, normally a riser cracking zone, operating at catalytic cracking conditions by contacting the hydrocarbon feed with a source of hot, regenerated cracking catalyst to produce an effluent comprising cracked products and spent catalyst containing coke and strippable hydrocarbons;
  • the effluent from the cracking zone is discharged and separated, normally in one or more cyclones, into a vapor phase rich in cracked products and a solids rich phase comprising the spent catalyst;
  • the vapor phase is removed as product and fractionated in the FCC main column and its associated side columns to form liquid cracking products including gasoline and light cycle oil;
  • At least a portion of the light cycle oil is recycled to a secondary catalytic cracking zone, preferably a second riser, wherein it contacts a fresh catalyst or freshly regenerated catalyst;
  • the spent catalyst is stripped, usually with steam, to remove occluded hydrocarbons from the catalyst, after which the stripped catalyst is oxidatively regenerated to produce hot, regenerated catalyst which is then recycled to the cracking zone for cracking further quantities of feed.
  • the feed to the FCC process will typically be a high boiling feed of mineral oil origin, normally with an initial boiling point of at least about 55O 0 F (29O 0 C) and in most cases above about 600 0 F (315 0 C). Most refinery cut points for FCC feed will be at least about 650 0 F (345°C). The end point will vary, depending on the exact character of the feed or on the operating characteristics of the refinery.
  • FCC feeds can include virgin feeds such as gas oils, e.g. heavy or light atmospheric gas oil, heavy or light vacuum gas oil as well as cracked feeds such as light coker gas oil, heavy coker gas oil as well as resid (non-distillable) material. Hydrotreated feeds may also be used, for example, hydrotreated gas oils, especially hydrotreated heavy gas oil. When utilizing the process of the present invention, it may be possible to dispense with initial hydrotreatment where its objective is to reduce sulfur although improvements in crackability will still be achieved.
  • FCC reactor riser top temperature conditions can be controlled in the range of about 900 to about 1050 0 F (about 482 to 565°C), preferably about 925° to about 1050 0 F (about 496 to 565°C) with typical operation at about 1000 0 F (about 54O 0 C).
  • most preferred FCC reactor riser top temperatures conditions for use of the present invention are on the lower end of these temperatures, preferably in the range of about 930 to about 97O 0 F (510 to 520 0 C).
  • Typical regenerated catalyst temperatures are in the range of about 1250 to about 1350 0 F (about 675 to 73O 0 C).
  • Catalystoil ratios from about 1: 1 to 20:1, preferably from 3:1 to 6:1, are typical. Pressures in the FCC reactor riser are normally of about atmospheric to about 350 kPag (50 psig) are preferred. These values are, however, subject to variation as discussed below if the generation of hindered DBTs in the process is to be mitigated according to the present process.
  • the feed is usually preheated to about 350° to 700 0 F (175 to 370 0 C), though operation with feed preheat outside of this range is possible.
  • the liquid cracking products from the FCC process typically include cracked naphtha fractions (light gasoline and heavy gasoline) boiling up to about 430 0 F (22O 0 C), and a full-range LCO (or “distillate") fraction typically boiling in the range of about 395 to about 75O 0 F (200 to about 400 0 C).
  • a undercut LCO fraction (such as the recycled LCO fraction herein) may also be drawn directly from an FCC fractionator or may be further separated from a full-range LCO fraction.
  • substantially as used in the disclosure herein, it is meant that at least 80 wt% of the designated fraction boils in the range of temperatures designated.
  • the cracking component of the FCC catalyst which is present to effect the desired cracking reactions and the production of lower boiling cracking products is typically based on a faujasite zeolite active cracking component, which is conventionally zeolite Y in one of its forms such as calcined rare-earth exchanged type Y zeolite (CREY), the preparation of which is disclosed in U.S. Pat. No. 3,402,996, ultrastable type Y zeolite (USY) as disclosed in U.S. Pat. No. 3,293,192, as well as various partially exchanged type Y zeolites as disclosed in U.S. Pat. Nos. 3,607,043 and 3,676,368.
  • CREY calcined rare-earth exchanged type Y zeolite
  • Cracking catalysts such as these are widely available in large quantities from various commercial suppliers.
  • the active cracking component is routinely combined with a matrix material such as silica and/or alumina as well as a clay in order to provide the desired mechanical characteristics (attrition resistance etc.) as well as activity control for the very active zeolite component or components.
  • the particle size of the cracking catalyst is typically in the range of 10 to 100 microns for effective fluidization. If separate particle additive catalysts are used, they are normally selected to have a particle size and density comparable to that of the cracking catalyst so as to prevent component segregation during the cracking cycle.
  • transalkylation onto the dibenzothiophenes is favored in the present processes by the use of catalysts with a large unit cell size in the zeolite component and a high matrix activity and/or high metals content.
  • the preferred cracking catalysts are those that have a low unit cell size. Unit cell sizes below 2.427 nm and lower, below 2.425 nm, are therefore preferred for the zeolite component.
  • low matrix activity and low metals content may also be favorable for low transalkylation activity, with matrix activity as measured by matrix surface area not more than 40 m 2 /gram and preferably not more than 35 or 30 m 2 /gram, in order to minimize the extent of transalkylation onto the unhindered DBT molecules present in the feed.
  • matrix activity as measured by matrix surface area not more than 40 m 2 /gram and preferably not more than 35 or 30 m 2 /gram
  • the effect of transalkylation onto the DBTs present in the feed is mitigated by a reversal of the process by which they form; in other words, the conditions under which the undesired
  • transalkylation takes place are replicated although optionally modified to favor transalkylation away from the hindered alkyl DBTs. If the hindered DBTs are given another chance to react, the equilibrium may be shifted and the amount of hindered sulfur in the resulting LCO changed.
  • the translkylation is carried out in a smaller, secondary riser of a dual-riser FCC unit which is fed directly with the LCO fraction and, also, with fresh or freshly regenerated cracking catalyst from the regenerator of the unit.
  • Figure 2 herein shows that the mono-alkyl and di-alkyl substituted DBTs are found principally in the highest boiling fractions of the LCO; it is these fractions, therefore, that are the most likely to benefit from any treatment which reduces the level of hindered alkyl DBTs.
  • the fractions representing the highest boiling 60% of the LCO fraction with boiling points substantially in the range of about 500 to about 75O 0 F (260 to 41O 0 C), and more preferably with boiling points substantially in the range of about 520 to about 75O 0 F (271 to 410 0 C) are the ones preferably treated in the present processing scheme.
  • the recycled fraction of LCO has boiling points substantially in the range of about 520 to about 68O 0 F (271 to 360 0 C). This is explained further in Example 2 herein.
  • the optimal final boiling point for the recycled LCO fraction can be determined empirically as a function of base FCC feed composition, catalyst selection, and operating conditions.
  • transalkylation does not require the high temperatures required for the actual cracking, lower temperatures are preferred in the secondary riser, favoring the transalkylation away from the hindered DBTs to the other species present in the selected LCO fraction.
  • the temperature should be adequate to vaporize the recycled LCO fraction in the riser.
  • a preferred target range being about 930 to about 97O 0 F (499 to 521 0 C), preferably about 950 to about 970 0 F (510 to 52 TC), for the present invention, this is especially desired conditions for the second reaction zone (or second riser) to wherein the LCO is recycled.
  • Riser top temperature can be controlled by appropriate selection of catalyst:oil ratio and regenerated catalyst temperature although the catalyst temperature required for cracking in the main riser will be the predominant consideration in selecting the temperature of the regenerated catalyst.
  • a relatively low catalystoil ratio coupled with a high regenerated catalyst temperature may be required to ensure feed vaporization with enough cooling in the riser to attain the desired riser top temperature.
  • Resort may also be made to riser quench to control riser top temperature, using quench media such as cycle oil, naphtha, distillate, waste oil.
  • Riser quench enables the mix zone temperature to be increased, typically by about 25 to about 50 0 F (15 to 3O 0 C) while still retaining the desired riser top temperature.
  • catalyst choice has been found to affect the efficacy of the alkyl transfer reactions.
  • Catalysts in which the zeolite component has high unit cell size tend to promote transalkylation onto the DBTs.
  • High matrix activity of a catalyst is also believed to be associated with high transalkylation activity.
  • catalysts with relatively lower unit cell size are less active for transalkylation and lower matrix activity may also be found to be associated with reduced
  • transalkylation activity This implies that if transalkylation of the DBT molecules is to be minimized to the extent feasible during the initial cracking reactions, a catalyst with low transalkylation activity would be the catalyst of choice (low unit cell size possibly coupled with low matrix activity).
  • transalkylation activity should desirably be maximized by using a catalyst of high unit cell size coupled potentially with high matrix activity. Because the FCC unit has to be operated with only one circulating catalyst however, a fundamental tension is established as it is not possible to accommodate both requirements simultaneously in one catalyst. A compromise catalyst candidate may therefore be the best choice although a final selection will be made on an empirical basis, taking into account the feed composition, product slate desired, unit characteristics and catalyst availability.
  • zeolite unit cell size of at least 2.425 nm, preferably at least 2.428 or even 2.430 nm have been found to confer good transalkylation activity with very notable results achieved with a zeolite unit cell size of at least 2.44 nm.
  • Embodiments of the present invention incorporating catalysts with a high activity matrix of at least 40 or even 50 ni 2 /gram surface area is also preferred.
  • VGO vacuum gas oil
  • Dibenzothiophene was added to the feed in amounts of 1%, 3% and 5%, to give nominal total sulfur contents of 1.15 wt.%, 1.47 wt%, and 1.77 wt.%,
  • each feed sample was run in the unit 4 to 5 times under the same conditions using ReduxionTM ECat (BASF) catalyst. Unless otherwise stated, each run in the unit was conducted at 99O 0 F (approximately 530 0 C) and a cat/oil ratio of 6.
  • the total sulfur content in the in total liquid product recovered from the process was obtained while the sample was still cold. The presence of the added DBT did not appreciably affect the conversion under the selected reaction conditions.
  • a positive number indicates that the 4,6 Dimethyl DBT structure is converted while a negative number indicates that the 4,6 Dimethyl DBT species is generated in the unit.
  • This study shows that the most promising way to reduce the concentration of 4,6 dimethyl DBT species is to recycle the 70-90% cut of this LCO (nominal boiling point 325-360 0 C, 620-680 0 F) at a low temperature.
  • the low cracking temperature enhances the transalkylation chemistry and ensures that the amount of dry gas and coke make is minimized.

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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)
  • Crystallography & Structural Chemistry (AREA)
  • Production Of Liquid Hydrocarbon Mixture For Refining Petroleum (AREA)

Abstract

La teneur en dibenzothiophènes alkylés à empêchement stérique dans la fraction intermédiaire d'un distillat issu d'un processus de craquage catalytique en lit fluidisé est réduite grâce au fractionnement des produits liquides craqués issus du processus en une fraction de naphte craqué et en une fraction de gasoil léger de craquage catalytique contenant des dibenzothiophènes alkylés. Une partie de la fraction de gasoil léger de craquage catalytique contenant les dibenzothiophènes alkylés, en général la fraction dont le point d'ébullition se situe essentiellement dans un intervalle de température d'environ 260 à 410 °C, est recyclée en direction d'une colonne secondaire en vue de la transalkylation, par d'autres espèces, des dibenzothiophènes alkylés formés durant les réactions initiales de craquage, dans la charge d'alimentation en hydrocarbures lourds.
PCT/US2010/046569 2009-08-28 2010-08-25 Réduction des dibenzothiophènes à empêchement stérique dans un distillat de craquage catalytique en lit fluidisé issu d'une unité de craquage catalytique en lit fluidisé à double zone réactionnelle WO2011025801A1 (fr)

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US12/584,005 2009-08-28

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Citations (6)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US4310489A (en) * 1980-08-14 1982-01-12 Standard Oil Company (Indiana) Apparatus for the catalytic cracking of hydrocarbons
US5897768A (en) * 1997-02-28 1999-04-27 Exxon Research And Engineering Co. Desulfurization process for removal of refractory organosulfur heterocycles from petroleum streams
US20040065618A1 (en) * 2001-02-16 2004-04-08 Ghaham Walter Ketley Purification process
US20050189260A1 (en) * 1998-12-28 2005-09-01 Chester Arthur W. Gasoline sulfur reduction in fluid catalytic cracking
US20060231459A1 (en) * 2005-03-28 2006-10-19 Swan George A Iii FCC process combining molecular separation with staged conversion
US20070227351A1 (en) * 2004-04-23 2007-10-04 Massachusetts Institute Of Technology Mesostructured Zeolitic Materials, and Methods of Making and Using the Same

Patent Citations (6)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US4310489A (en) * 1980-08-14 1982-01-12 Standard Oil Company (Indiana) Apparatus for the catalytic cracking of hydrocarbons
US5897768A (en) * 1997-02-28 1999-04-27 Exxon Research And Engineering Co. Desulfurization process for removal of refractory organosulfur heterocycles from petroleum streams
US20050189260A1 (en) * 1998-12-28 2005-09-01 Chester Arthur W. Gasoline sulfur reduction in fluid catalytic cracking
US20040065618A1 (en) * 2001-02-16 2004-04-08 Ghaham Walter Ketley Purification process
US20070227351A1 (en) * 2004-04-23 2007-10-04 Massachusetts Institute Of Technology Mesostructured Zeolitic Materials, and Methods of Making and Using the Same
US20060231459A1 (en) * 2005-03-28 2006-10-19 Swan George A Iii FCC process combining molecular separation with staged conversion

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