WO2011025801A1

WO2011025801A1 - Reduction of hindered dibenzothiophenes in fcc distillate from a dual reaction zone fcc unit

Info

Publication number: WO2011025801A1
Application number: PCT/US2010/046569
Authority: WO
Inventors: Stacey E. Siporin; George A. Swan, Iii; Bruce R. Cook; Steven S. Lowenthal; Michael A. Hayes; Michael W. Bedell; Steve Colgrove
Original assignee: Exxonmobil Research And Engineering Company
Priority date: 2009-08-28
Filing date: 2010-08-25
Publication date: 2011-03-03

Abstract

The level of sterically hindered alkylated dibenzothiophenes in the middle distillate fraction from an FCC process is reduced by fractionating the cracked liquid products of the process to form a cracked naphtha fraction and a light cycle oil fraction that contains alkylated dibenzothiophenes. A portion of the light cycle oil fraction that contains alkylated dibenzothiophenes, typically the fraction boiling substantially in the range from about 500 to about 750°F (260 to 410°C), is recycled to a secondary reaction riser to transalkylate the alkylated dibenzothiophenes formed during the initial cracking reactions with other species in the heavy hydrocarbon feed.

Description

REDUCTION OF HINDERED DIBENZOTHIOPHENES IN FCC

DISTILLATE FROM A DUAL REACTION ZONE FCC UNIT

FIELD OF THE INVENTION

[0001] This invention relates to a process for producing low sulfur distillates such as low sulfur road diesel fuel.

BACKGROUND OF THE INVENTION

[0002] Environmental concerns are expected to lead to decreases in the permissible levels of sulfur in hydrocarbon fuels. While reduction in the maximum sulfur level of road diesel oils from about 0.3 weight percent to 0.05 weight percent were implemented in the 1990s, further significant reductions have since come into effect. In the European Union, the Euro IV standard specifying a maximum of 50 wppm (0.005%) of sulfur in diesel fuel for most highway vehicles has applied since 2005; ultra-low sulfur diesel with a maximum of 10 wppm of sulfur was required to be available from 2005 and was, in fact, widely available in 2008. A final target is the 2009 Euro V fuel standard for the final reduction of sulfur to 10 wppm, which is also expected for most non-highway applications.

[0003] In the United States, the Environmental Protection Administration ("EPA") has required most on-highway diesel fuel sold at retail locations in the United States to conform to the Ultra Low Sulfur Diesel ("ULSD") standard of 15 wppm since 2006 except for rural Alaska which will transition all diesel to ULSD in 2010. Non-road diesel fuel, required to conform to 500 wppm sulfur in 2007, will be further limited to conform to ULSD sulfur specifications in 2010 and railroad locomotive and marine diesel fuel will also change to conform with ULSD sulfur specifications in 2012. As of December 1, 2014 all highway, non- road, locomotive and marine diesel fuel produced and imported into the United States will be required to conform with the ULSD specifications.

[0004] The current allowable sulfur content for ULSD in the United States (15 wppm) is much lower than the previous U.S. on-highway standard for low sulfur diesel ("LSD") of 500 wppm. The reduced sulfur content not only reduces emissions of sulfur compounds but also allows advanced emission control systems to be fitted that would otherwise be poisoned by these compounds. These systems can greatly reduce emissions of oxides of nitrogen and particulate matter and according to EPA estimates, emissions of nitrogen oxide will be reduced by 2.3 million metric tonnes (2.6 million short tons) each year and soot or particulate matter will be reduced by 100,000 metric tonnes (110,000 short tons) a year with the adoption of the new standards.

[0005] 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 Merox™ 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

dibenzothiophenes ("DBTs"). Conventional hydrodesulfurization processes are capable of removing sulfur compounds, especially the lower molecular weight materials including the aliphatic sulfur materials, thiophenes and

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. For example, 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

Processes, Gates et al. McGraw Hill, pages 407 and 408.

[0006] Hydrogenative removal of the dibenzothiophenes requires high hydrogen partial pressures and circulation rates, low space velocity and high temperature, implying a significant increase in the capacity of the hydrogen circulation system, an increase in the reactor bed size, an increase in operating pressure, a decrease in cycle length or any combination of these. The higher severity operation can also increase cracking and, therefore, light gas production. Conventional hydroprocessing of the fractions which find their way into the light diesel products such as road diesel is therefore economically unattractive as a complete solution.

[0007] One of the fractions which is conventionally used as a blend component for road diesel is light cycle oil ("LCO") which is produced in large quantities in the fluid catalytic cracking ("FCC") units commonly used for gasoline production. Experience has shown, however, that hydrodesulftirization of full-range LCO requires high severity conditions for achieving sulfur levels as low as 500 ppm let alone the far lower levels required by ULSD or Euro V.

These difficulties are, moreover, accentuated by the fact that the other sulfur-and nitrogen-containing impurities in the feed react earlier in the reactor, producing ammonia and hydrogen sulfide, which further inhibit the removal of the dibenzothiophenes.

[0008] Improvements in hydroprocessing techniques such as those described in U.S. Patent No. 5,409,599 and 5,730,860, were developed in response to the previous sulfur limitations and improved management of refinery operations both in FCC units and in the hydroprocessing of middle distillates has achieved worthwhile reductions but at some cost. Refiners may, for example, choose to undercut their LCO to remove the hindered DBTs from the molecules that need to be hydrotreated. Undercutting is a practice of lowering the end boiling point (or "cut point") of a hydrocarbon fraction and thus sends more valuable, lower boiling point hydrocarbons from the LCO to the FCC bottoms resulting in a significant loss in revenue.

[0009] Other refiners choose to use very high pressure hydrogen to desulfurize the hindered DBTs in the FCC products. This process is also very costly absent significant improvements in the activity of diesel

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. Some attention has been given to adding components such as Zn and Ni as additives to FCC catalysts to trap the sulfur in the hydrocarbons. In such instances, the sulfur would be sent to the regenerator as ZnS instead of as organic sulfur in the product streams. The sulfur would then exit the regenerator as SO_x where it would need to be further treated. 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.

[0010] Managing sulfur reduction in middle distillate fuels has created an incentive to develop improved methods for removing sulfur compounds, especially the refractory alkylated dibenzothiophenes, in ways which are economical as well as effective.

SUMMARY OF THE INVENTION

[0011] We have found that a significant amount of alkylation of the DBT molecules occurs in the FCC unit. That is that a non-hindered DBT molecule entering the FCC unit is likely to undergo transalkylation with alkyl fragments cracked off other molecules and leave the unit as a sterically hindered DBT which becomes more difficult to hydrotreat. We have discovered herein a way to mitigate the effect of alkylation of DBT molecules to hindered DBTs in the FCC unit so that the proportion of hindered, refractory DBT molecules in the FCC product fractions can be reduced without incurring the cost of either rejecting useful LCO to the lower-value heavy oil products or without the use of high severity hydroprocessing of the DBT compounds.

[0012] In an embodiment of 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. However, 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. [0013] Accordingly, therefore, the level of hindered alkylated

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 75O⁰F (260 to 400°C)_Λ and more preferably, the fraction boiling substantially in the range from about 520 to about 75O⁰F (271 to 400⁰C)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. In a most preferred embodiment of the present invention, the recycled portion of the LCO fraction boils substantially in the range from about 520 to about 68O⁰F (271 to 36O⁰C). By substantially it is meant that at least 80 wt% of the designated fraction boils in the range of temperatures designated.

[0014] 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.

FIGURES

[0015] 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. [0016] FIGURE 2 is a graph showing the sulfur speciation of a typical light cycle oil.

[0017] FIGURE 3 is a graph showing the sulfur speciation of cracking products of a vacuum gas oil ("VGO") with dibenzothiophene added in three concentrations.

DETAILED DESCRIPTION

FCC Unit Configuration

[0018] Figure 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. 5,248,411 and 5,376,339 and alternative systems such as the UOP VSS™ system, are well established and widely known. 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).

[0019] 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 Process

[0020] The predominate commercially utilized catalytic cracking process for gasoline production currently in use is the fluid catalytic cracking ("FCC") process. Apart from the use of the LCO recycle, the manner of operating the FCC process of the present invention will remain essentially unchanged. Thus, 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.

[0021] Briefly, 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. The significant steps in the cyclic process of the invention are:

(i) 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;

(ii) 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; (iii) 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;

(iv) 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; and

(iv) 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.

[0022] 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 55O⁰F (29O⁰C) and in most cases above about 600⁰F (315⁰C). Most refinery cut points for FCC feed will be at least about 650⁰F (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.

[0023] FCC reactor riser top temperature conditions can be controlled in the range of about 900 to about 1050⁰F (about 482 to 565°C), preferably about 925° to about 1050⁰F (about 496 to 565°C) with typical operation at about 1000⁰F (about 54O⁰C). However, 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 97O⁰F (510 to 520⁰C). Typical regenerated catalyst temperatures are in the range of about 1250 to about 1350⁰F (about 675 to 73O⁰C). 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 700⁰F (175 to 370⁰C), though operation with feed preheat outside of this range is possible.

[0024] The liquid cracking products from the FCC process typically include cracked naphtha fractions (light gasoline and heavy gasoline) boiling up to about 430⁰F (22O⁰C), and a full-range LCO (or "distillate") fraction typically boiling in the range of about 395 to about 75O⁰F (200 to about 400⁰C). 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. By use of the term "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.

FCC Catalyst

[0025] 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. 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. Generation of Hindered Dibenzothiophenes

[0026] It has been discovered that a significant portion of the sterically hindered dibenzothiophenes generated during the FCC process arise from the alkylation of simpler dibenzothiophenes ("DBT's) present in the feed with alkyl fragments liberated during the cracking; generation from other sulfur species such as alkylthiophenes, sulfides, mercaptans has not been found to be significant. It was further discovered that the choice of cracking catalyst had an effect on the extent to which the transalkylation took place during the cracking. As a general proposition, 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. For this reason, when the basic process objective is to minimize the generation of hindered DBTs in the initial cracking step, 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. It is also believed that 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 m²/gram and preferably not more than 35 or 30 m²/gram, in order to minimize the extent of transalkylation onto the unhindered DBT molecules present in the feed. As discussed below, however, a strategy of minimizing the generation of hindered DBTs by some reversal of the undesired alkylation is favored by the use of a catalyst that increases the degree of transalkylation, so establishing a tension in the final choice of catalyst. Hindered DBT Reduction Strategy

[0027] According to the present invention, 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.

[0028] Speciation of a typical light cycle oil fraction (boiling substantially in the range of from about 350 to 75O⁰F) from a refinery FCC unit has shown that the hindered DBTs are concentrated in the heaviest portions of the LCO. Figure 2 shows the sulfur distribution, by Sulfur Simdist, in an illustrative refinery LCO. Table 1.1 below shows the boiling point distribution of the LCO and Table 1.2. shows the sulfur distribution, by Sulfur Simdist, of the same refinery LCO by boiling point.

Table 1.1

Sim Dist. Boiling Point of LCO Cuts

Table 1.2

Total Sulfur of LCO Cuts by Boiling Point

[0029] 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 75O⁰F (260 to 41O⁰C), and more preferably with boiling points substantially in the range of about 520 to about 75O⁰F (271 to 410⁰C) are the ones preferably treated in the present processing scheme. It has been found, however, that it in a most preferred embodiment of the present invention, it may not be preferable to recycle the highest boiling fractions of the LCO with boiling points above about 680⁰F (360⁰C) since these contain long chain alkyl substituents which on cracking may generate C₁ and C₂ alkyl fragments which may react with DBT cores and so re-form the hindered DBTs such as 4,6 dimethyl DBT and 4 ethyl DBT which should be removed.

Therefore, in the most preferable embodiment of the present invention the recycled fraction of LCO has boiling points substantially in the range of about 520 to about 68O⁰F (271 to 360⁰C). 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.

[0030] Because 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 97O⁰F (499 to 521⁰C), preferably about 950 to about 970⁰F (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.

[0031] 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 50⁰F (15 to 3O⁰C) while still retaining the desired riser top temperature.

[0032] As noted above, 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 (as measured, for example, by matrix surface area) is also believed to be associated with high transalkylation activity. Conversely, 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). On the other hand, when the LCO fraction herein is recycled, 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.

[0033] Although the trends relating the zeolite cell size and matrix activity of the catalyst to transalkylation activity are not firmly fixed, 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 ni²/gram surface area is also preferred.

[0034] The use of freshly regenerated or fresh, steamed catalyst from the unit regenerator is particularly favorable for promoting the desired new equilibrium between the DBTs as this catalyst has higher activity than the partly spent catalyst which is used in the processes of the two related application where the recycled LCO fractions is re-equilibrated in the basic cracking cycle no earlier in the cycle than at the riser top.

Example 1

[0035] Model compound spiking experiments in a laboratory scale FCC unit were carried out to better understand the FCC sulfur chemistry. The feed used was, throughout, a vacuum gas oil ("VGO") containing 0.99 wt. pet. sulfur.

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.%,

respectively.

[0036] Each feed sample was run in the unit 4 to 5 times under the same conditions using Reduxion™ ECat (BASF) catalyst. Unless otherwise stated, each run in the unit was conducted at 99O⁰F (approximately 530⁰C) and a cat/oil ratio of 6. The sulfur distribution in the total liquid product ("TLP") was determined on the basis of simulated distillation with the hindered alkyl DBTs (C₁ DBT = monosubstituted DBT, C₂ DBT= monosubstituted DBT) showing longer retention times in the Simdist chromatogram than the unhindered (CO) DBT. 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.

[0037] The results shown in Figure 3 indicated that dibenzothiophenes ("DBT"s) in the feed were a major contributor to distillate range sulfur. A significant amount of alkylation occurred in the unit to form hindered DBTs but no substantial cracking of DBTs was observed nor was there any evidence of ring growth to form coke. Formation of DBTs or hindered DBTs was not observed with other experiments in which other sulfides and thiophenes were added to the feed. For example, no olefin and H₂S combination to form thiophenes was evident from runs where octyl sulfide was added to the feed.

Example 2

[0038] Individual cuts of the LCO reported in Tables 1.1 and 1.2 above were fed directly into the laboratory scale FCC unit, with a catalyst of unit cell size 2.429 nm and matrix activity of 30 m²/g. The reaction was run at two different temperatures, namely 530⁰C and 505⁰C (99O⁰F and 94O⁰F). A summary of the results is shown in Table 2.

Table 2

4,6 Dimethyl DBT conversion in FCC

[0039] 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 325-360⁰C, 620-680⁰F) 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. Recycling the heaviest cut (90-100%) of the LCO results in 4,6 dimethyl DBT generation due to the long alkyl chains (>=C₂) cracking to form C₁ and C₂ species such as 4,6 dimethyl DBT and 4 ethyl DBT.

Claims

WHAT IS CLAIMED IS:

1. A fluid catalytic cracking process comprising catalytically cracking a heavy hydrocarbon feed in a fluid catalytic cracking (FCC) unit comprising the steps of:

(a) contacting the hydrocarbon feed with a heated circulating catalyst in a first reaction zone of a fluid catalytic cracking process;

(b) producing cracked hydrocarbon product fractions including a cracked naphtha fraction and a light cycle oil (LCO) fraction containing alkylated dibenzothiophenes; and

(c) passing at least a portion of the light cycle oil fraction containing alkylated dibenzothiophenes to a second reaction zone of the fluidized catalytic process, wherein the at least 80 wt% of the portion of the light cycle oil fraction that is recycled boils in the range from about 500 to about 75O⁰F (260 to 41O⁰C), thereby transalkylating at least a portion of the alkylated dibenzothiophenes in the presence of circulating fluid catalytic cracking catalyst;

wherein the circulating catalyst comprises a faujasitic zeolite with a unit cell size of at least 2.425 nra; and the level of hindered alkylated

dibenzothiophenes in the cracked hydrocarbon product fractions is reduced.

2. The process of claim 1, wherein at least 80 wt% of the light cycle oil fraction containing alkylated dibenzothiophenes boils in the range from about 520 to about 75O⁰F (271 to 400⁰C).

3. The process of claim 1, wherein the FCC unit is a unit with a primary reaction riser and a secondary reaction riser and at least a portion of the LCO fraction containing alkylated dibenzothiophenes is recycled to the secondary reaction riser.

4. The process of claim 3, wherein the portion of the LCO fraction containing alkylated dibenzothiophenes recycled to the secondary reaction riser is contacted with freshly regenerated cracking catalyst from an FCC regenerator.

5. The process of claim 4, wherein the riser top temperature of the secondary reaction riser is from about 930 to about 970⁰F (499 to 52 TC).

6. The process of claim 5, wherein the faujasitic zeolite has a unit cell size of at least 2.430 nm.

7. The process of claim 6, wherein the surface area of the faujasitic zeolite is at least 40 m²/gram.

8. The process of claim 7, wherein the cracked hydrocarbon product fractions are comprised of a full-range LCO fraction that has a boiling range substantially from about 395 to about 570⁰F (200 to 300⁰C) and the level of hindered alkylated dibenzothiophenes in the full-range LCO fraction is reduced.

9. The process of claim 8, wherein the faujasitic zeolite has a unit cell size of at least 2.440 nm.

10. A fluid catalytic cracking process in which a heavy hydrocarbon feed containing organosulfur compounds is catalytically cracked to lighter products by contact of the feed in a cyclic catalyst recirculation cracking process with a circulating inventory of a fluidizable catalytic cracking catalyst comprising catalyst particles, comprising the steps of:

(a) catalytically cracking the hydrocarbon feed in a first reaction zone of a fluid catalytic cracking process operating at catalytic cracking conditions by contacting the hydrocarbon feed with a source of hot, regenerated cracking catalyst to produce an effluent comprising cracked hydrocarbon products and spent catalyst which contains coke and strippable hydrocarbons;

(b) separating the effluent from the first reaction zone into cracked hydrocarbon products and spent catalyst;

(c) fractionating the cracked hydrocarbons products to form a light cycle oil fraction comprising alkylated dibenzothiophenes;

(d) stripping the spent catalyst in a stripper section to remove occluded hydrocarbons from the spent catalyst to produce a stripped catalyst;

(e) oxidatively regenerating the stripped catalyst in a regenerator to produce a regenerated catalyst;

(f) recycling the regenerated cracking catalyst to the first reaction zone and a second reaction zone; and

(g) recycling a portion of the light cycle oil fraction comprising alkylated dibenzothiophenes to the second reaction zone, wherein the at least 80 wt% of the portion of the light cycle oil fraction that is recycled boils in the range from about 500 to about 75O⁰F (260 to 41O⁰C), thereby transalkylating at least a portion of the alkylated dibenzothiophenes in the recycled light cycle oil fraction in the presence of the regenerated cracking catalyst;

wherein the regenerated cracking catalyst comprises a faujasitic zeolite with a unit cell size of at least 2.425 nm; and the level of hindered alkylated dibenzothiophenes in the cracked hydrocarbon products is reduced.

11. The process of claim 10, wherein at least 80 wt% of the recycled LCO fraction comprising alkylated dibenzothiophenes boils in the range from about 520 to about 68O⁰F (271 to 36O⁰C).

12. The process of claim 10_? wherein the FCC unit is a unit with a primary reaction riser and a secondary reaction riser and at least a portion of the LCO fraction comprising alkylated dibenzothiophenes is recycled to the second reaction zone which is located in the secondary reaction riser.

13. The process of claim 12, wherein the riser top temperature of the secondary reaction riser is from about 930 to about 970⁰F (499 to 521⁰C).

14. The process of claim 13, wherein the surface area of the faujasitic zeolite is at least 40 m /gram.

15. The process of claim 14, wherein the faujasitic zeolite has a unit cell size of the zeolite is at least 2.430 nm.

16. The process of claim 15, wherein at least a portion of the LCO fraction comprising alkylated dibenzothiophenes recycled to the secondary reaction zone is contacted with a fresh cracking catalyst.

17. The process of claim 16, wherein the cracked product fractions are comprised of a full-range LCO fraction that has a boiling range substantially from about 395 to about 570⁰F (200 to 300⁰C) and the level of hindered alkylated dibenzothiophenes in the full -range LCO fraction is reduced.