GB2411851A - Process for upgrading FCC product with additional reactor - Google Patents

Abstract

A process is disclosed for taking a cut from an FCC reactor 10 product and reacting it in a separate reactor 50 to upgrade the product quality. Cracking or reformulating reactions in the separate reactor give reductions in olefins and reformulating hydrogen-transfer reactions convert undesirable olefins to isoparaffins and aromatics without reducing octane value. Catalyst particles from the FCC reactor may be cycled to the separate reactor. This process has also been found to substantially diminish concentrations of nitrogen and sulfur compounds fed to the separate reactor.

Description

2411 851 "PROCESS FOR UPGRADING FCC PRODUCT WITH ADDITIONAL REACTOR"

BACKGROUND OF THE INVENTION

1] This invention relates generally to processes for the fluidized catalytic cracking (FCC) of heavy hydrocarbon streams. More specifically, this invention relates generally to processes for upgrading catalytically cracked hydrocarbon feeds in a discrete reactor vessel.

DESCRIPTION OF THE PRIOR ART

2] The FCC process is carried out by contacting the starting material whether it be vacuum gas oil, reduced crude, or another source of relatively high boiling hydrocarbons with a catalyst made up of finely divided or particulate solid material. The catalyst is transported in a lo fluid-like manner by passing gas or vapor through it at sufficient velocity to produce a desired regime of fluid transport. Contact of the oil with the fluidized material catalyzes the cracking reaction. The cracking reaction deposits coke on the catalyst. Coke is comprised of hydrogen and carbon and can include other materials in trace quantities such as sulfur and metals that enter the process with the starting material. Coke interferes with the catalytic activity of the catalyst by blocking active sites on the catalyst surface where the cracking reactions take place. Catalyst is traditionally transferred from a stripper, that removes adsorbed hydrocarbons and gases from catalyst, to a regenerator for purposes of removing the coke by oxidation with an oxygen containing gas. An inventory of catalyst having a reduced coke content, relative to the catalyst in the stripper, hereinafter referred to as regenerated catalyst, is collected for return to the reaction zone. Oxidizing the coke from the catalyst surface releases a large amount of heat, a portion of which escapes the regenerator with gaseous products of coke oxidation generally referred to as flue gas. The balance of the heat leaves the regenerator with the regenerated catalyst. The fluidized catalyst is continuously circulated from the reaction zone to the regeneration zone and then again to the reaction zone. The fluidized catalyst, as well as providing a catalytic function, acts as a vehicle for the transfer of heat from zone to zone. Catalyst exiting the reaction zone is spoken of as being spent, i.e., partially deactivated by the deposition of coke upon the catalyst.

The FCC processes, as well as separation devices used therein are fully described in US-A 5,584,985 B 1 and US-A-4,792,437 B 1, the contents of which are hereby incorporated by reference. Specific details of the various contact zones, regeneration zones, and stripping zones along with arrangements for conveying the catalyst between the various zones are well known to those skilled in the art.

3] The FCC reactor cracks gas oil or heavier feeds into a broad range of products.

Cracked vapors from the FCC unit enter a separation zone, typically in the form of a main column, that provides a gas stream, a gasoline cut, light cycle oil (LCO) and clarified oil (CO) which includes heavy cycle oil (HCO) components. The gasoline cut may include light, medium and heavy gasoline components. A major component of the heavy gasoline fraction comprises condensed single ring aromatics. A major component of LCO is condensed bicyclic ring aromatics.

4] Subjecting product fractions to additional reactions is useful for upgrading product quality. The recracking of heavy product fractions from the initially cracked FCC product is one example. Typically, in recracking, uncracked effluent from a first riser of an FCC reactor is recontacted with catalyst at a second location to cleave larger molecules down into more useful smaller molecules. For example, US-A-4,051,013 B 1 discloses cracking both gasoline-range feed and gas oil feed in the same riser at different elevations. US-A-3,161,582 B1, US-A 5,176,815 B1 and US-A-5,310,477 B1 all disclose cracking a primary hydrocarbon feedin a riser of an FCC unit and cracking a secondary hydrocarbon feed in a reactor into which the riser exits. As a result, both cracked products mix in the reactor, to some extent, which could negate the incremental upgrade resulting from cracking the secondary hydrocarbon feed, particularly if it is a fraction of the cracked primary hydrocarbon feed.

5] US-A-3,928,172 Bl teaches an FCC unit with a secondary dense fluidized catalyst bed in a separate reactor. Gas oil is cracked in a riser of the FCC unit with unregenerated catalyst from the separate dense fluidized catalyst bed. A heavy naphtha fraction of the cracked gas oil, boiling between 127 and 2320C (260 and 450 F), from the riser is recracked in the separate reactor over regenerated catalyst. Apparently, the benefit of cracking lower boiling fractions was not explored, presumably because the octane rating of the lower boiling fraction was sufficiently high or because it was predicted not to be effective. The data in the patent indicates that nominal, if any, reformulation reactions occur in the separate reactor because little, if any, new aromatics are produced.

6] In gasoline production, many governmental entities are restricting the concentration of olefins allowed in the gasoline pool. Reducing olefin concentration without also reducing value is difficult because higher olefin concentrations typically promote higher Research Octane Numbers (RON) and Motor Octane Numbers (MON), but the latter to a lesser extent. Octane value or Road Octane Number is the average of RON and MON. Merely saturating olefins typically yields normal paraffins which typically have low octane value. Additionally, saturation requires the addition of hydrogen, which is expensive and in some regions, difficult to obtain.

7] Feedstocks for FCC units typically include sulfur and nitrogen. During FCC operation, the sulfur and nitrogen are converted primarily to hydrogen sulfide and ammonia, which are easily removed, but are also converted to organic sulfurs, mercaptans and nitrogen oxides. Stricter environmental limits on sulfur and nitrogen compound emissions along with lower sulfur specifications for fuel products have raised interest in the need to remove nitrogen and sulfur compounds from FCC gasoline. As demands for cleaner fuels and use of high sulfur and high nitrogen feedstocks increase, the need for sulfur and nitrogen removal from FCC gasoline will become even greater.

t0008] It is an object of the present invention to provide a method for enhancing the quality of product from an FCC unit. It is a further object of the present invention to reduce the olefinicity of product from an FCC unit without substantially reducing the octane rating of the product and without the addition of hydrogen. It is an even further object of the present invention to reduce the concentration of sulfur and nitrogen compounds in an FCC product.

BRIEF SUMMARY OF THE INVENTION

9] It has now been discovered that a separate reactor can be used to either reformulate or crack a product fraction from an FCC unit to reduce its olefinicity and maintain or boost its octane rating without the separate addition of hydrogen. If the separate reactor is incorporated lo into an FCC unit, catalyst can be circulated between the FCC reactor and the separate reactor.

Additionally, it has been further found that higher boiling point fractions from an FCC unit can be hydrotreated and sent to a separate reactor, if incorporated in the FCC unit using catalyst cycled through the FCC unit, to crack FCC product fractions down to lower boiling point useful hydrocarbon components. Furthermore, reacting fractions of FCC product in a separate reactor has been found effective in substantially reducing sulfur and nitrogen compounds in the fraction.

0] Additional objects, embodiment and details of this invention can be obtained from the following detailed description of the invention.

BRIEF DESCRIPTION OF THE DRAWINGS

1] FIG. 1 is a sectional, elevational, schematical view of an FCC unit incorporating a main column and a secondary reactor in accordance with the present invention.

2] FIG. 2 is a sectional, elevational, schematical view of an alternative embodiment of the present invention.

3] FIG. 3 is a sectional, elevational, schematical view of a further embodiment of the present invention.

DETAILED DESCRIPTION OF THE INVENTION

4] The present invention may be described with reference to four components: an FCC reactor IO, a regenerator 50, a secondary reactor 80, 80', 80" and a main column 100. Although many configurations of the present invention are possible, three specific embodiments are presented herein by way of example. All other possible embodiments for carrying out the present invention are considered within the scope of the present invention. For example, the secondary reactor 80, 80', 80" and/or the main column 100 need not be incorporated into an FCC unit as illustrated in FIGS. 1-3 but may stand alone.

5] In the embodiment of the present invention in FIG. 1, the FCC reactor 10 comprises a conduit in the form of a reactor riser 12 that extends upwardly through a lower portion of a reactor vessel 14 as in a typical FCC arrangement. The central conduit or reactor riser 12 preferably has a vertical orientation within the reactor vessel 14 and may extend upwardly through the bottom of the reactor vessel 14 or downwardly from the top of the reactor vessel 14.

The reactor riser 12 terminates in a separation vessel 16 at swirl arms 18. A hydrocarbon feed stream is fed to the riser at a nozzle 20 which is contacted and vaporized by hot regenerated catalyst fluidized by a gas such as steam from a nozzle 22. The catalyst cracks the hydrocarbon feed stream and a mixture of catalyst particles and gaseous cracked hydrocarbons exit the swirl arms 18 into the separation vessel 16. Tangential discharge of gases and catalyst from the swirl arms 18 produces a swirling helical Notion about the interior of the separation vessel 16, causing heavier catalyst particles to fall into a dense catalyst bed 24 and a mixture of gaseous cracked hydrocarbons and entrained catalyst particles to travel up a gas recovery conduit 26 and enter into cyclones 28. In the cyclones 28, centripetal force imparted to the mixture induces the heavier entrained catalyst particles to fall through diplegs 30 of the cyclone 28 and to the bottom of the separation vessel 16 into a dense catalyst bed 32. The gases in the cyclones 28 more easily 2s change direction and begin an upward spiral with the gases ultimately exiting the cyclones 28 through outlet pipes 34. Cracked gases leave the reactor vessel 14 though an outlet conduit 36.

The cracked gases are optionally subjected to a further separation (not shown) to further remove any light loading of catalyst particles and are sent via a line 98 to fractionation in the main column 100 which will be described later with reference to all of FIGS. 1-3. Catalyst particles in the dense catalyst bed 32 enter the separation vessel 16 through windows 38 where they join catalyst particles in the dense catalyst bed 24 in a stripping section 40 of the separation vessel 16.

The catalyst particles are stripped of entrained cracked vapors over bames 42 with a stripping medium such as steam entering from at least one nozzle 44. The stripped cracked vapors travel up to the gas recovery conduit 26 where they are processed with other cracked product vapors.

to [0016] Stripped catalyst from the stripping section 40 of the FCC reactor 10 travels through a first stripped catalyst pipe 46 regulated by a control valve 48 and into the regenerator 50 at a lower chamber 52. In the lower chamber 52, stripped catalyst is subjected to hot oxygen containing gas such as air from a distributor 54. Coke is burned from the catalyst and as the catalyst is heated, it ascends upwardly in the lower chamber 52 and is distributed into an upper chamber 55 of the regenerator through a distributor 56. Regenerated catalyst collects in a dense catalyst bed 58 whereas entrained catalyst is removed from regenerator effluent gases in cyclones and 62. Flue gas exits the cyclone 62 through an outlet pipe 64 to exit the regenerator through an outlet 66. Regenerated catalyst from the dense catalyst bed 58 travels through a regenerated catalyst pipe 68 regulated by a control valve 70 into the reactor riser 12 where it is fluidized and contacted with fresh feed. Stripped catalyst also exits the stripping section 40 through a second stripped catalyst pipe 72 regulated by a control valve 74 into a dense catalyst bed 82 in the secondary reactor 80. The degree to which the control valve 74 is opened can be automatically controlled to obtain the temperature desired in the secondary reactor 80. For example, if higher temperature is desired in the secondary reactor 80, more of the relatively hot catalyst can be permitted to pass through the control valve 74 to add heat to the secondary reactor 80. The secondary reactor 80 is preferably a fluidized bed. However, a riser reactor or other reactor configuration may be suitable. A partition defines a hopper section 8 l of the secondary reactor 80. Catalyst in the dense catalyst bed 82 that falls into the hopper section 81 is fluidized by steam or some other fluidizing media through a distributor 84 and is stripped of entrained gases s over baffles 83. A desired cut of hydrocarbon feed from the FCC reactor 10 and fractionated in the main column 100 is fed to a secondary reactor 80. The feed to the secondary reactor 80 from the main column lOO is fed through a distributor 86 where it is contacted with catalyst in the dense catalyst bed 82. The distributor 86 distributes feed in such a way as to fluidize the dense catalyst bed 82. Cyclones 88 and 90 remove entrained catalyst from a gaseous product which leaves the secondary reactor 80 through a conduit 92. Catalyst leaves the secondary reactor 80 after being stripped in the hopper section 8 l through a pipe 76 regulated by a control valve 78.

The degree to which the control valve 78 is opened can be automatically controlled to obtain the level desired in the secondary reactor 80. The level of the catalyst in the secondary reactor 80 determines the weight hourly space velocity (WHSV) of reactants through the secondary reactor Is 80. For example, if a greater WHSV is desired, the control valve 78 would be opened relatively more to reduce the level of catalyst in the dense catalyst bed 82.

7] FIG. 2 is an alternative embodiment of the present invention in which regenerated catalyst is fed to the secondary reactor 80'. In FIG. 2, the elements of the FCC reactor 10 and the regenerator 50 have generally the same configuration as in FIG. 1. Elements in FIG. 2 with different configurations from FIG. 1, such as in the secondary reactor 80', will be distinguished by adding a " ' " symbol to the reference numeral. Hydrocarbon feed processed in the FCC reactor 10 is recovered at the outlet conduit 36 and is canted by the line 98 to be fractionated in the main column 100, perhaps after interim processing, to obtain a desired cut to be fed to the secondary reactor 80'. The feed to the secondary reactor 80' is fed by a fluidizing nozzle 85 to be 2s contacted in a riser 86' with regenerated catalyst from a regenerated catalyst pipe 68' regulated by a control valve 70'. Both feed and catalyst are distributed by the riser 86' into a dense catalyst bed 82' which is fluidized by the feed from the riser 86'. Products exit the secondary reactor 80' out a conduit 92' after entrained catalyst is removed in cyclones 88' and 90'. A partition defines a hopper section 81' of the secondary reactor 80'. Catalyst from the dense catalyst bed 82' in the secondary reactor 80' that falls into the hopper section 81' is fluidized with a medium such as steam from a distributor 84' and is stripped of entrained product gases over dames 83'. Stripped catalyst passes through a pipe 76' regulated by a control valve 78' to the reactor riser 12 where it contacts the primary hydrocarbon feed stream injected by the nozzle 20. Stripped catalyst from the stripping section 40 of the FCC reactor 10 passes through a stripped catalyst pipe 46' regulated by a control valve 48' into the lower chamber 52 of the regenerator 50 where coke deposits are burned from catalyst by means of a hot oxygen- containing gas such as air.

Regenerated catalyst from the upper chamber 55 passes through the regenerated catalyst pipe 68' and is regulated by the control valve 70' before it enters the riser 86' of the secondary reactor 80'. All other elements in FIG. 2 have generally the same function as in FIG. 1.

8] FIG. 3 shows another embodiment of an FCC unit utilizing a secondary reactor 80" which receives catalyst from and returns catalyst to the regenerator 50. Again, because the FCC reactor 10 and the regenerator 50 are both very similar to those depicted in FIG. 1, all of their elements in both drawings will retain the same reference numerals. However, those elements in FIG. 3 that differ from the corresponding elements in FIG. 1 will be distinguished by adding a " " " symbol to the reference numeral. Primary hydrocarbon feed is fed to the reactor riser 12 by means of the nozzle 20. The primary feed is contacted with regenerated catalyst and cracked to yield product that is withdrawn from the FCC reactor 10 via the outlet conduit 36. Catalyst separated from the cracked product is stripped m the stripping section 40 and passed through a stripped catalyst pipe 46" regulated by a control valve 48" into the lower chamber 52 of the 2s regenerator 50. Regenerated catalyst from the upper chamber 55 of the regenerator 50 is distributed to the reactor riser 12 through a first regenerated catalyst pipe 68" regulated by a control valve 70" where it contacts fresh primary feed and is also distributed through a second regenerated catalyst pipe 72" regulated by a control valve 74" to the secondary reactor 80". The gaseous vapor effluent from the FCC reactor 10 is carried from the outlet conduit 36 through the line 98, perhaps to further processing and then to the main column 100 to be fractionated. A desired fraction is fed to the secondary reactor 80" through a distributor 86" which fluidizes a dense catalyst bed 82" with a medium such as steam. The feed contacts regenerated catalyst in the dense catalyst bed 82". A partition defines a hopper section 81" in the secondary reactor 80".

Catalyst from the dense catalyst bed 82" of the secondary reactor 80" that falls into the hopper section 81" is fluidized by steam of some other fluidizing media through a distributor 84" and is stripped of entrained gases over baffles 83". Stripped catalyst passes through a pipe 76" regulated by a control valve 78" to the regenerator 50. The product from the secondary reaction is recovered through cyclones 88" and 90" which remove entrained catalyst and send the catalyst back to the dense catalyst bed 82". A conduit 92" carries gaseous product to further processing which could consist of heating and fractionating.

9] The secondary reactor 80. 80'. 80" may stand alone instead of being incorporated into an FCC unit. If the secondary reactor 80, 80', 80" stands alone, the preferred feed will be a cut of product from an FCC unit.

0] In reference to all of FIGS. 1-3, the cracked product stream in the line 98 from the FCC reactor 10, relatively free of catalyst particles and including the stripping fluid, exits the reactor vessel 14 through the outlet conduit 36. The cracked product stream in the line 98 may be subjected to additional treatment to remove fine catalyst particles or to further prepare the stream prior to fractionation. The l ne 98 transfers the product stream containing the cracked product to a fractionator in the form of the main column 100. A variety of products are withdrawn from the main column 100. In this case, the main column 100 recovers an overhead stream of light products comprising unstabilized gasoline and lighter gases. A line 102 transfers the overhead stream through a condenser 104 and a cooler 106 before it enters a receiver 108. A line 110 withdraws a light off-gas stream from the receiver 108. A bottom liquid stream of light gasoline leaves the receiver 108 via a line 112 which may have to undergo further treatment to stabilize the light gasoline. The main column 100 also provides a heavy gasoline stream, an LCO stream and an HCO stream through lines 120, 122 and 124, respectively. Parts of the streams in the lines 120, 122 and 124 are all circulated through heat exchangers 126, 128 and 130 and reflux loops 132, 134 and 136, respectively, to remove heat from the main column 100. Streams of heavy gasoline, LCO and HCO are transported from the main column 100 through respective lines 140, 142 and 144. A CO fraction may be recovered from the bottom of the main column via a line 146. Part of the CO fraction is recycled through a reboiler 148 and returned to the main column 100 through a line 150. The CO stream is removed from the main column 100 via a line 152.

1] The light gasoline or light naphtha fraction preferably has an initial boiling point (IBP) below 127 C (260 F) in the Cs range; i.e., 35 C (95 F), and an end point (EP) at a temperature greater than or equal to 127 C (260 F). The boiling points for these fractions are determined using the procedure known as ASTM D86-82. The heavy gasoline or heavy naphtha fraction has an IBP at or above 127 C (260 F) and an EP at a temperature above 200 C (392 F), preferably between 204 and 221 C (400 and 430 F), particularly at 216 C (420 F). The LCO stream has an IBP at about the EP temperature of the heavy gasoline and an EP in a range of 260 to 371 C (500 to 700 F) and preferably 288 C (550 F). The HCO stream has an IBP of the EP temperature of the LCO stream and an EP in a range of 371 to 427 C (700 to 800 F), and preferably 399 C (750 F). The CO stream has an IBP of the EP temperature of the HCO stream and includes everything boiling at a higher temperature. One or more of each of these streams or 2s other cuts from the main column 100 are sent to the secondary reactor 80, 80', 80" to be contacted with the catalyst therein. In one embodiment, a stream such as the hne 142 which carries LCO may be hydrotreated in a hydrotreating reactor 154 before it is sent to the secondary reactor 80, 80', 80" for cracking. Other streams from the main column 100 could be hydrotreated before entering the secondary reactor 80, 80', 80".

2] In the secondary reactor 80, 80', 80", the predominant reaction may be cracking in which a hydrocarbon molecule is broken into two smaller hydrocarbon molecules, so that the number of carbon atoms in each molecule diminishes. Alternatively, the predominant reaction in the secondary reactor 80, 80', 80" may be a hydrogen-transfer reaction such as reformulation or isomerization in which the structures of the molecules are changed but the number of carbon 0 atoms in each molecule does not change. In determining which type of reaction, cracking or hydrogen transfer, predominates over the other, reactions involving compounds with 5 to 8 carbons may be the most relevant because they include most of the olefins which can either crack or reform.

3] Olefins, naphthenes and cyclo-olefins are reformulated into paraffins, aromatics and some naphthenes as shown in formulas (1), (2), (3) and (4).

3 Cn H2n + Cm H9m 3 Cn H2n+2 + Cm H2m-6 (1) olefins + naphthene paraffins + aromatic 4 Cn H2n 3 Cn H2n+2 + Cn H2n-6 (2) olefins paraffins + aromatic Cm H2m 2 + 2 Cn H2n Cm H2m-6 + 2 Cn H2n+2 (3) cyclo-olefins + olefins aromatic + paraffins Cn H2n + H2 Cn H2n+2 (4) olefins + hydrogen paraffins Olefins have a higher octane value than their paraffinic counterpart. Hence, the conversion of olefins to paraffins typically degrades octane value. When the olefins cyclitize to become aromatics as shown in formulas (1) and (2) and when cyclo-olefins aromaticize to yield aromatics as in formula (3), they donate much hydrogen. Other olefins pick up the hydrogen to become paraffins as shown in formula (4). In the present invention using the secondary reactor 80, 80', 80", normal olefins and iso-olefins predominantly reformulate to isoparaffins which carry a higher octane rating than normal paraffins. Additionally, aromatics also boost the octane rating of the product. Because the isoparaffins and aromatics have a high octane rating, the hydrogen transfer reformulation in the secondary reactor 80, 80', 80" maintains the high octane ratings despite the typical octane rating decline that accompanies conversion of olefins to paraffins. Accordingly, the hydrogen-transfer reactions in the secondary reactor 80, 80', 80" which yield more isoparaffins and aromatics are superior to a process which saturates the olefins into normal paraffins. Advantageously, the hydrogen transfer reactions are performed without the addition of hydrogen, which can be expensive and difficult to obtain.

4] Production of aromatics is a gauge for the degree of hydrogen transfer that occurs in the reaction. When conditions are set to promote hydrogen transfer reactions in the secondary reactor 80, 80', 80", a net yield increase in aromatics of 5% on a fresh feed basis is typical and at least a 40% increase is easily attainable.

5] The reaction in the secondary reactor 80, 80', 80" is preferably conducted with the same catalyst circulated through the regenerator 50 and the FCC reactor 10. Of course, if a secondary reactor 80, 80', 80" stands alone without incorporation into an FCC unit, the catalyst in the secondary reactor need not be circulated through an FCC unit. If hydrogentransfer reactions are intended to predominate over cracking reactions in the secondary reactor, the WHSV will typically range from 0.1 to 5 hrl. If cracking reactions are to predominate over hydrogen-transfer reactions, the WHSV will typically range from 5 to 50 hrl. Additionally, the conditions in a hydrogen-transfer reaction are less severe, with temperatures in the range of 399 to 510 C (750 to 950 F) than in a cracking reaction with temperatures in the range of 482 to 649 C (900 to 1200 F).

6] An additional advantage of the hydrogen transfer reaction in the secondary reactor 80, 80', 80" is that it is endothermic. Hence, the spent catalyst which contacts the hydrocarbon stream in the dense catalyst bed 82, 182, 282 is cooled before it is sent back to the reactor riser 12 of the FCC reactor 10 or the regenerator 50. Consequently, heat will be removed from the whole system which permits use of a greater catalystto-oil ratio in the reactor riser 12, resulting in higher conversion in the FCC reactor 10.

7] The reformulation of the fraction from the main column 100 by hydrogen transfer in the secondary reactor 80, 80', 80" reduces the concentrations of organic sulfur and nitrogen compounds in the products. The reaction of the gasoline fraction in the secondary reactor 80, 80', 80" can lower sulfur concentration in the reactor products by as much as 80 wt-% and nitrogen concentration in the products by as much as 98 wt-%. Hence, the products from the secondary reactor 80, 80', 80" will contain low concentrations of sulfur and nitrogen compounds. Leftover sulfur and nitrogen compounds can be removed from the product by hydrotreating and taken off in the overhead of a finishing distillation column if necessary to meet

specifications.

8] Typically, the catalyst circulation rate through the reactor riser 12 and the input of feed and any lift gas that enters the riser will produce a flowing density of between 48 and 320 kg/m3 (3 and 20 Ibs/ft3) and an average velocity of 3 to 31 rn/sec (10 to 100 ft/sec) for the catalyst and gaseous mixture. In the FCC reactor 10, catalyst will usually contact the hydrocarbons in a catalyst to oil ratio in a range of from 3 to 8, and more preferably in a range of from 4 to 6. The length of the reactor riser 12 will usually be set to provide a residence time of between 0.5 to 10 seconds at these average flow velocity conditions. Other reaction conditions in the reactor riser 12 usually include a temperature of from 468 to 566 C (875 to 1050 F).

9] This invention can employ a wide range of commonly used FCCcatalysts. These catalyst compositions include high activity crystalline alumina silicate or zeolite containing catalysts. Zeolite catalysts are preferred because of their higher intrinsic activity and their higher resistance to the deactivating effects of high temperature exposure to steam and exposure to the metals contained in most feedstocks. Zeolites are usually dispersed in a porous inorganic carrier material such as silica, aluminum, or zirconium. These catalyst compositions may have a zeolite content of 30% or more. Zeoltes including high silica-to-alumina compositions such as LZ-210 and ZSM-5 type materials are preferred when lighter products are desired. Another particularly useful type of FCC catalysts comprises silicon substituted aluminas. As disclosed in US-A 5,080,778 B1, the zeolite or silicon enhanced alumina catalysts compositions may include intercalated clays, also generally known as pillared clays. The preferred catalysts for the present invention include USY zeolites. When hydrogen-transfer reactions are desired to predominate over cracking reactions in the secondary reactor 80, 80', 80", high rare earth content Y zeolites are preferred. The term "high rare earth content" denotes greater than 2.0 wt-% rare earth oxide on the zeolite portion of the catalyst. High rare earth content Y zeolites such as USY zeolite may have as much as 4 wt-% rare earth. The high rare earth content promotes hydrogen transfer by increasing adjacent acid site density on the catalyst. Strongly acidic catalyst sites on the catalyst promote cracking. Y zeolites with low rare earth content can still effectively promote hydrogen transfer but with longer reactor residence times. When cracking reactions are desired to predominate over hydrogen transfer reactions in the secondary reactor 80, 80', 80", low rare earth Y zeolite catalysts are preferred which have a rare earth oxide content of 2.0 wt-% or less.

Additives, such as sulfur-reducing additives, may be added to the catalyst. It is anticipated that such additives may experience enhanced effectiveness in the secondary reactor for longer residence times.

0] Feeds suitable for processing by this invention include conventional FCC feedstocks or higher boiling hydrocarbon feeds. The most common of the conventional feedstocks is a Is vacuum gas oil which is typically a hydrocarbon material having a boiling range of from 343 to 552 C (650 to 1025 F) and is prepared by vacuum fractionation of atmospheric residue. Such fractions are generally low in coke precursors and heavy metals which can deactivate the catalyst.

1] When LCO is the feed to the secondary reactor 80, 80', 80", a portion of the LCO fraction will typically pass through the hydrotreating reactor 154 and be transported through a line 156 to the secondary reactor 80, 80', 80" in which J-crackng occurs. When operating in the LCO mode of this invention, the LCO cut carries bicyclic aromatic compounds into the secondary reactor 80, 80', 80" which cannot be cracked unless they are pretreated. These bicyclic compounds include indenes, biphenyls and naphthalenes which are refractory to cracking under lo the conditions in the reactor riser 12. In the J-cracking process, one of the rings of the bicyclic hydrocarbons are saturated. The saturated ring is then cracked in the secondary reactor 80, 80', 80" and cleaved from the aromatic ring as shown in exemplary formulas (5) and (6). CH3

3 jig + 6 H2 3 JCH3 (5) CH3 CH3 2 + 2 H3C -\/CH3 CH3 Abyss 'CH3 (6) [0032] In formula (5), one of the rings of dimethyl naphthalene is saturated to make dimethyl tetrahydronaphthalenes. In formula (6), the saturated ring of two dimethyl tetrahydronaphthalenes are cracked and accept hydrogen donated from a ring of another dimethyl tetrahydronaphthalene that aromaticizes. The cracked rings yield toluene and isobutane.

3] Suitable methods for carrying out J-cracking are further described in US-A-3,479,279 B 1 and US-A-3,356,609 B 1 which are incorporated herein by reference. The J-cracking process eliminates about two-thirds of the high boiling aromatics from an LCO cut bringing the effluent from the secondary reactor 80, 80', 80" into the gasoline boiling range. The LCO fraction can pass through the hydrotreatng reactor 154 as a separate stream or together with another fraction from the main column 100.

4] The hydrotreatment of the fraction in the hydrotreating reactor 154 takes place at low severity conditions to avoid the saturation of the single ring aromatic compounds in the gasoline fraction. In the method of this invention, up to 100% of the fraction may be hydrotreated.

Hydrotreating is carried out in the presence of a nickel-molybdenum or cobalt-molybdenum catalyst and relatively mild hydrotreating conditions including a temperature of 316 to 371 C (600 to 700 F), a liquid hourly space velocity (LHSV) of from 0.2 to 2 hr1 and a pressure of 3447 to 10342 kPa (500 to 1500 psi").

EXAMPLES

EXAMPLE 1

5] A fraction of gasoline from an FCC reactor effluent having the properties in Table I was subjected to coked USY zeolite catalyst with 1 to 1.5 wt-% rare earth in a reactor at the conditions in Table I. The reaction yielded a product with the properties in Table I.

TABLE I

FEED PROPERTIES

IBP, C ( F) 121 (250) Aromatics, wt-% 61.8 Olefins, wt-% 14.2

_

Paraffins/Naphthenes, wt-% 24

_

RON 93.3 MON 81.9

REACTOR CONDITIONS

WHSV, hr1 1 Reaction Temperature, C ( F) 454 (850) Catalyst-to-Oil Ratio 6.0 Pressure, kPa (psi") 69 (10)

PRODUCT PROPERTIES

C2-, wt-% 0.6 C3, wt-% 1.2

_

C4, wt-% 2.0

_

Cs+ / 232 C (450 F), wt-% 89.4 LCO, wt-% 4.7 CO, wt-% _ 2.1 Gasoline RON 95.8 Gasoline MON 84 Aromatics, wt-% 70 Olefins, wt-% 1 Paraffins/Naphthenes, wt-% 29 In this example, the olefin concentration dropped from 14.2% to 1 wt-% as a result of the secondary reaction. Whereas, the aromatics concentration increased from 61.8 to 70 wt-%.

Additionally, both the RON and the MON increased. The relatively small concentrations of C4 and smaller hydrocarbons reveal that cracking reactions were minor compared to the s reformulating, hydrogen transfer reactions indicated by the increase in aromatics.

EXAMPLE 2

6] A separate study was performed to determine the effect on product properties of four sets of operating conditions on full range FCC gasoline as shown in Table II.

TABLE II

FEED PROPERTES

IBP, C ( F) 35 (95)

_

Paraffins, wt-% 27 Olefins, wt-% 51

__

Naphthenes, wt-% 6 Aromatics, wt-% 14 C4, wt-% 2.3 Feed Boiling Over 221 C (430 F), wt-% 1.3 _. _

PROCESS CONDITIONS A B C D

Reaction Temperature, C ( F) 399 (750) 399 (750) 454 (850) 482 (900) Catalyst-to-Oil Ratio 3 5 5.1 5.1 PRODUCT YELDS, wt-% C2- 0.06 0.13 0. 43 0.60 C3 0.82 1.22 2.85 4.16 C4 3.5 4.53 6.75 8.35 Cs+ / 220 C (429 F) 91.3 86.4 83.1 80.0 LCO 2.5 3.69 3.28 2.87 CO 0.2 1.5 1.4 1.9

__

Coke 1.6 2.5 2.2 2.1 Gasoline Recovery 94.9 90.0 86.7 83.6 Paraffins 42 47 48 44 Olefins 31 21 18 _ 13 Naphthenes 8 8 7 7 Aromatics 21 23 27 36 As the temperature is increased, the gasoline recovery diminished while the aromatics concentration increased and the olefins concentration decreased. Additionally, cracking as indicated by the amount of C4 and lower carbon number concentration increases as the reaction temperature andlor catalyst-to-oil ratio increases. Accordingly, the reaction conditions can be tailored to obtain a desired product quality.

EXAMPLE 3

7] The feed in the next set of experiments had the properties given in Table III.

TABLE III

Paraffins wt-% 28.1 Olefins, wt-% 50.4 Naphthenes, wt-% 5.9 Aromatics, wt% 14.4 C l Non-Aromatics, wt-% 1.32 RON 91.0 MON 79.3 Road Octane Number 85.2 Sulfur, ppm 136 Nitrogen, ppm 46 Ca, wt-% 2.3

_

221 C (430 F) plus, wt-% 1.3 IBP, C ( F) 35 (95) T10 51 (123) T30 67 (153) T50 88 (190) T70 118 (244) T90 152 (306) EP, C ( F) 179 (354) The foregoing feed was reacted under three different sets of conditions with corresponding product yields and quality given in Table IV.

TABLE IV Run

PROCESS CONDITIONS A B C

Reactor Temperature, C ( F) 427 (800) 454 (850) 482 (900) Catalyst-toOil Ratio 6.5 6.1 5.9 Hydrocarbon Partial Pressure, kPa (psia) 117 (17.0) 114 (16.5) 122 (17.7) System Pressure, kPa (psi") 278 (40.3) 276 (40.0) 273 (39.6) LHSV, hr1 4.6 4.6 4.6 PRODUCT YIELDS, wt-% Dry Gas 0.4 0.7 1.1

_

c3's 1.6 2.4 3.4 C4's 6.1 7.8 9.4 Cs+ Gasoline 85.5 83.0 80.0 Paraffins 53.3 54.7 52.3 Olefins 13.8 12.4 12.3 Naphthenes 8.1 5.5 6.2

_

Aromatics 24.8 27.4 29.2 Sulfur, ppm 69 62 68 - . ....

Nitrogen, ppm 1 2 4 RON 87.4 88.4 90.4 MON 80.5 81.5 81.8 Road Octane Number 84.0 85.0 _ 86.1 The foregoing qualities and yields pertaining to the Cs+ gasoline have been adjusted to reflect the fact that C4's were present in the feed which did not participate in the reaction and would not be present in the feed to the secondary reactor. Moreover, the data indicates that not much cracking occurred in the reaction because relatively small quantities of C4- material is generated.

The process also reduces the olefin concentration while increasing the paraffin and aromatics concentration, all without substantial change in the Road Octane Number.

8] Table V gives the breakdown of the product composition from foregoing Run B by carbon number and compound type. The number that is not in parentheses in Table V is the weight percentage of that compound in the feed. Whereas, the number in parentheses is the weight percentage of the compound in the product. With regard to Table V, aromatics with nine or more carbon numbers are grouped together. Therefore, the numbers given for carbon numbers and 11 in the "Total" column include only non- aromatic Clo's and C1 1's. The minimal changes in total concentration of each carbon number fraction, especially in the Cs-C range shows that reformulating hydrogen transfers are predominant over cracking reactions under this set of conditions. Moreover, the large increase in isoparaffins compared to the moderate increase in paraffins greatly offsets the octane value debit resulting from olefin reduction.

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Claims (10)

1. CLAIMS: 1. A process for converting a hydrocarbon feed stream comprising:

   cracking a preliminary cracking feed stream with catalyst particles in a cracking reactor (12) to produce a cracked product, said catalyst particles having a composition s including crystalline alumina silicate or zeolite; separating said cracked product from said catalyst particles in a separator vessel (16) to obtain a cracked product stream; recovering at least a portion of said cracked product stream to be a reformulation feed stream; lo cycling catalyst particles that had resided in said cracking reactor to a reformulating reactor that is discrete from said separator vessel; passing said reformulation feed stream including saturated and olefinic hydrocarbons with carbon numbers of 5-8 to said reformulating reactor (80, 80', 80"); reformulating said reformulation feed stream in said reformulating reactor (80, 80', 80") to produce a reformulated product stream, said reformulating proceeding at conditions that promote at least a 5% net yield increase in aromatics on a fresh reformulation feed basis indicating the occurrence of hydrogen transfer reactions; and recovering said reformulated product stream.
2. 2. The process of claim 1 further including isolating said reformulated product stream from said cracked product stream.
3. 3. The process of claim 1 wherein said catalyst particles in said reformulating reactor have the same composition as the catalyst particles in said cracking reactor.
4. 4. The process of claim 1 wherein a greater proportion of hydrocarbons with carbon numbers of 5-8 undergo hydrogen transfer reaction than cracking reaction.
5. S. The process of claim 1 wherein olefins in said reformulation feed stream convert to isoparaffins in the reformulating reactor (80, 80', 80") .
6. 6. The process of claim 1 wherein the concentration of sulfur compounds in the reformulated product stream is less than its concentration in the reformulation feed stream.
7. 7. The process of claim 1 wherein the concentration of nitrogen compounds in the reformulated product stream is less than its concentration in the reformulation feed stream.
8. 8. The process of claim 1 wherein the reformulation feed stream has an initial boiling point below 200 C.
9. 9. The process of claimlwherein the reformulation feed stream is a naphtha stream having an initial boiling point below 127 C.
10. 10. The process of claimlwherein the reformulation feed stream is an oil stream from said cracked product stream having an initial boiling point of at least 200 C.