Abstract:
A process is disclosed for taking a cut from an FCC reactor product and reacting it in a separate reactor 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:
BACKGROUND OF THE INVENTION  
         [0001]    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  
         [0002]    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 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 U.S. Pat. No. 5,584,985 B1 and U.S. Pat. No. 4,792,437 B1, 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.  
           [0003]    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.  
           [0004]    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, U.S. Pat. No. 4,051,013 B1 discloses cracking both gasoline-range feed and gas oil feed in the same riser at different elevations. U.S. Pat. No. 3,161,582 B1, U.S. Pat. No. 5,176,815 B1 and U.S. Pat. No. 5,310,477 B1 all disclose cracking a primary hydrocarbon feed in 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.  
           [0005]    FCC units employing two risers are known. U.S. Pat. No. 5,198,590 B1, U.S. Pat. No. 4,402,913 B1, U.S. Pat. No. 4,310,489 B1, U.S. Pat. No. 4,297,203 B1, U.S. Pat. No. 3,799,864 B1, U.S. Pat. No. 3,748,251 B1, U.S. Pat. No. 3,714,024 B1 and WO 00/40672 disclose two riser FCC units in which feeds are predominantly cracked in both risers In these patents, both risers communicate with the same recovery conduit and/or reactor permitting commingling of gaseous products. In U.S. Pat. No. 5,730,859 B1, all of the effluent from one riser is fed to the other riser, without first undergoing a product separation. U.S. Pat. No. 4,172,812 B1 teaches recracking all or a part of cracked product from a riser of an FCC unit over a catalyst having a composition that is different from the catalyst composition in the riser.  
           [0006]    In U.S. Pat. No. 5,944,982 B1, although both risers terminate in the same reactor vessel, gaseous products from each riser are isolated from the other. This patent also discusses the method of cracking a LCO fraction of a cracked product containing very refractory, bicyclic aromatic components. Bicyclic aromatics are very difficult to crack and boil only at high temperatures. The presence of high boiling point bicyclic aromatics can cause a gasoline pool to exceed maximum volatility standards. By a process called J-cracking, LCO is hydrotreated to partially saturate the bicyclic aromatic hydrocarbons such as naphthalene to produce tetralin. The tetralin is then cracked to make benzene, toluene, xylene and isoparaffins along with some naphthalene. The “J” in J-cracking is a measure of unsaturation of the hydrocarbons having the general formula:  
           C N H 2N−J .  
           [0007]    U.S. Pat. No. 3,928,172 B1 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 232° C. (2600 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.  
           [0008]    It is also known to subject cracked product from a riser of an FCC unit to a subsequent oligomerization reaction in a separate reactor. In oligomerization, smaller olefins are bonded together to make larger olefins of greater molecular weight. U.S. Pat. No. 5,009,851 B1 and U.S. Pat. No. 4,865,718 B1 disclose oligomerizing a fraction of cracked product from an FCC unit in a separate reactor.  
           [0009]    Cracking and oligomerization differ from reformulation in that the former involve decreases or increases in carbon numbers, respectively. Whereas, in reformulation, carbon numbers are not changed but hydrogen atoms are exchanged to alter the structure of the molecule and make it more valuable.  
           [0010]    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.  
           [0011]    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.  
           [0012]    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  
         [0013]    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 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.  
           [0014]    Accordingly, in one embodiment, the present invention relates to a process for converting a hydrocarbon feed stream comprising passing a reformulation feed stream including saturated and olefinic hydrocarbons with carbon numbers of 5-8 to a reformulating reactor. The reformulating reactor contains catalyst particles having a composition. The reformulation feed stream is reformulated in the reformulating reactor to produce a reformulated product stream. The reformulating proceeds 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. The reformulated product stream is then recovered.  
           [0015]    In another embodiment, the present invention relates to a process for converting a hydrocarbon feed stream comprising contacting the hydrocarbon feed stream with catalyst particles having a composition in a first reactor to produce a cracked product. The cracked product is separated from the catalyst particles in a vessel to obtain a cracked product stream. A naphtha stream is recovered from the cracked product stream. The naphtha stream has an initial boiling point below 127° C. (260° F.). The naphtha stream is contacted with catalyst particles having the composition in a second reactor to produce an upgraded product stream. The upgraded product stream is recovered and isolated from the cracked product stream.  
           [0016]    In a further embodiment, the present invention relates to a process for converting a hydrocarbon feed stream comprising contacting the hydrocarbon feed stream with catalyst particles having a composition in a first reactor to produce a cracked product. The cracked product is separated from the catalyst particles in a vessel to obtain a cracked product stream. An oil stream is recovered from the cracked product stream having an initial boiling point above about 200° C. (392° F.). Catalyst particles that had resided in the first reactor are cycled to a second reactor that is discrete from the vessel. The oil stream is contacted with catalyst particles in a second reactor to produce an upgraded product stream. The upgraded product stream is recovered and isolated from the cracked product stream.  
           [0017]    Additional objects, embodiment and details of this invention can be obtained from the following detailed description of the invention.  
       
    
    
     BRIEF DESCRIPTION OF THE DRAWINGS  
       [0018]    [0018]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.  
         [0019]    [0019]FIG. 2 is a sectional, elevational, schematical view of an alternative embodiment of the present invention  
         [0020]    [0020]FIG. 3 is a sectional, elevational, schematical view of a further embodiment of the present invention. 
     
    
     DETAILED DESCRIPTION OF THE INVENTION  
       [0021]    The present invention may be described with reference to four components: an FCC reactor  10 , 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 herdein 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.  
         [0022]    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 motion 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 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 baffles  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.  
         [0023]    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  60  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  81  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 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  100  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  81  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  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 .  
         [0024]    [0024]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 carried 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 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 baffles  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.  
         [0025]    [0025]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 in 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 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.  
         [0026]    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.  
         [0027]    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 line  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  100  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 .  
         [0028]    The light gasoline or light naphtha fraction preferably has an initial boiling point (IBP) below 127° C. (260° F.) in the C 5  range; i.e., about 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. (4000 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. (5000 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 3710 to 427° C. (7000 to 800° F.), and preferably about 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 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 line  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 ″.  
         [0029]    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 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.  
         [0030]    Olefins, naphthenes and cyclo-olefins are reformulated into paraffins, aromatics and some naphthenes as shown in formulas (1), (2), (3) and (4).  
         3C n H 2n +C m H 2m →3C n H 2n+2 +C m H 2m−6 olefins+naphthene→paraffins+aromatic  (1)  
         4C n H 2n →3C n H 2n+2 +C n H 2n−6 olefins+paraffins→aromatic  (2)  
         C m H 2m−2 +2C n H 2n →C m H 2m−6 +2 C n H 2n+2 cyclo-olefins+olefins→aromatic+paraffins  (3)  
         C n H 2n +H 2 →C n H 2n+2 olefins+hydrogen→paraffins  (4)  
         [0031]    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  
         [0032]    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.  
         [0033]    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 hydrogen-transfer reactions are intended to predominate over cracking reactions in the secondary reactor, the WHSV will typically range from 0.1 to 5 hr −1 . If cracking reactions are to predominate over hydrogen-transfer reactions, the WHSV will typically range from 5 to 50 hr −1 . Additionally, the conditions in a hydrogen-transfer reaction are less severe, with temperatures in the range of 3990 to 510° C. (750° to 950° F.) than in a cracking reaction with temperatures in the range of 482° to 649° C. (9000 to 1200° F.).  
         [0034]    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 catalyst-to-oil ratio in the reactor riser  12 , resulting in higher conversion in the FCC reactor  10 .  
         [0035]    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.  
         [0036]    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/m 3  (3 and 20 lbs/ft 3 ) and an average velocity of about 3 to 31 m/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. (8750 to 105° F.).  
         [0037]    This invention can employ a wide range of commonly used FCC catalysts. 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. Zeolites 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 U.S. Pat. No. 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 about 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.  
         [0038]    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 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.  
         [0039]    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-cracking 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 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).  
                         
 
         [0040]    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.  
         [0041]    Suitable methods for carrying out J-cracking are further described in U.S. Pat. No. 3,479,279 B1 and U.S. Pat. No. 3,356,609 B1 which are incorporated herein by reference. The J-cracking process eliminates about two-hirds 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 hydrotreating reactor  154  as a separate stream or together with another fraction from the main column  100 .  
         [0042]    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 3160 to 371° C. (6000 to 700° F.), a liquid hourly space velocity (LHSV) of from 0.2 to 2 hr −1  and a pressure of 3447 to 10342 kPa (500 to 1500 psig).  
         [0043]    The present invention can be operated in several ways, four of which are explained herein. In the first exemplary operation, higher proportions of LCO and LPG are obtained. The FCC reactor  10  is run at relatively low severity with a temperature between 482° and 521° C. (900° and 970° F.) and a short contact time of 1 to 3 seconds. The FCC reactor  10  will thus operate at low conversion to yield a high proportion of LCO, HCO and CO, some gasoline and some liquefied petroleum gas (LPG), all withdrawn from the main column  100 . If the feed to the FCC reactor  10  is highly paraffinic, all of the CO can be fed to the secondary reactor  80 ,  80 ′,  80 ″. However, if the feed is not highly paraffinic, only the HCO fraction should be fed to the secondary reactor. Fractions of LCO and LPG product can be recovered from the main column  100 . If gasoline is desired, it can be recovered from the main column and sent to the gasoline pool. If gasoline is not desired, it can be sent with CO and HCO or alone to the secondary reactor  80 ,  80 ′,  80 ″ which cracks the CO and HCO mixture at high severity temperatures such as 521° to 560° C. (9700 to 1040° F.) and preferably 549° C. (1020° F.) and at a space time of 1 to 10 hr −1 . LPG and LCO are then recovered from the secondary reactor which can be added to the fractions of LCO and LPG recovered from the main column  100 . A medium or smaller pore, shape selective zeolite additive such as ZSM-5 may be added to the catalyst to obtain greater yields of LPG in this operation. Because the secondary reactor is operated at high severity, FIG. 2 or  3  would be most appropriate for this operation because the hotter catalyst from the regenerator  50  can provide the necessary heat requirements.  
         [0044]    A second operation in which the present invention can be used to produce gasoline, LPG and benzene, toluene and xylene (BTX) gasoline. The FCC reactor  10  is run at a high severity temperature ranging from 521° to 560° C. (9700 to 1040° F.), preferably 549° C. (1020° F.) and a contact time of over 3 seconds. The high severity cracking operation gives a high conversion with gasoline, LPG, LCO and CO in the product stream. Gasoline and LPG are recovered from the main column  100  while LCO is fed from the main column  100  to the hydrotreating reactor  154  to saturate one of the bicyclic aromatic rings to prepare it for cracking. The hydrotreated LCO is then sent to the secondary reactor  80 ,  80 ′,  80 ″ operated at high severity temperatures of 521° to 560° C. (970° to 1040° F.) sufficient to J-crack it to obtain BTX gasoline which can be mixed with gasoline to upgrade gasoline product quality. The embodiments in FIG. 2 or  3  can be used for this exemplary operation.  
         [0045]    In a third exemplary operation, the desired product yields up to an 80% reduction in gasoline sulfur and nitrogen and possesses an olefin concentration as low as 1 wt- %. The primary reactor is run at a severity appropriate to obtain the desired conversion. Either a full range cut of gasoline having an IBP below 127° C. (260° F.) and an EP at or below 200° C. (392° F.) or a fraction thereof from the main column  100  is fed to the secondary reactor  80 ,  80 ′,  80 ″ which is run at 482° to 521° C. (900° to 970° F.). In the secondary reactor, the olefins reformulate via hydrogen transfer to isoparaffins and aromatics with minimal gasoline yield loss and an octane gain and without need of additional hydrogen. Moreover, sulfur levels are reduced by as much as 80 wt-% and nitrogen levels are reduced by as much as 98 wt-%. If necessary, the gasoline can then be hydrotreated to reduce sulfur and nitrogen compounds to even lower levels to meet specifications by converting them to hydrogen sulfide and ammonia, respectively, which can be removed in the light ends of a downstream gasoline fractionation unit (not shown) with minimal octane debit and consumption of hydrogen. This operation can be performed with any of the three embodiments in FIGS.  1 - 3  of the present invention.  
         [0046]    When the desired products are LCO and low olefinicity, moderate octane gasoline, a fourth exemplary operation may be used. The FCC reactor  10  is run at low severity at a temperature of 482° to 521° C. (9000 to 970° F.) and a contact time of 1 to 3 seconds. The low conversion operation yields high quantities of LCO, some gasoline and not much LPG. The LCO can be recovered from the main column  100 . The gasoline fraction can be fed to the secondary reactor at low severity 4820 to 521° C. (900 to 970° F.) and low WHSV, 0.1 to 5 hr −1 , so the gasoline reforms to convert olefins to aromatics and isoparaffins to upgrade the gasoline quality.  
       EXAMPLES  
     Example 1  
       [0047]    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, hr −1     1           Reaction Temperature, ° C. (° F.)   454 (850)           Catalyst-to-Oil Ratio   6.0           Pressure, kPa (psig)   69 (10)           PRODUCT PROPERTIES           C 2   − , wt-%   0.6           C 3 , wt-%   1.2           C 4 , wt-%   2.0           C 5   + /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                      
 
         [0048]    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 C 4  and smaller hydrocarbons reveal that cracking reactions were minor compared to the reformulating, hydrogen transfer reactions indicated by the increase in aromatics.  
       Example 2  
       [0049]    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 PROPERTIES           IBP, ° C. (° F.)   35 (95)       Paraffrns, wt-%   27       Olefins, wt-%   51       Naphthenes, wt-%   6       Aromatics, wt-%   14       C 4 , wt-%   2.3       Feed Boiling Over   1.3       221° C. (430° F.), wt-%            PROCESS   A   B   C   D       CONDITIONS       Reaction   399 (750)   399 (750)   454 (850)   482 (900)       Temperature, ° C. (° F.)       Catalyst-to-Oil Ratio   3   5   5.1   5.1       PRODUCT YIELDS,       wt-%       C 2   −     0.06   0.13   0.43   0.60       C 3     0.82   1.22   2.85   4.16       C 4     3.5   4.53   6.75   8.35       C 5   + /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                  
 
         [0050]    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 C 4  and lower carbon number concentration increases as the reaction temperature and/or catalyst-to-oil ratio increases. Accordingly, the reaction conditions can be tailored to obtain a desired product quality.  
       Example 3  
       [0051]    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 12  Non-Aromatics, wt-%   1.32           RON   91.0           MON   79.3           Road Octane Number   85.2           Sulfur, ppm   136           Nitrogen, ppm   46           C 4 , 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)                      
 
         [0052]    The foregoing feed was reacted under three different sets of conditions with corresponding product yields and quality given in Table IV.  
                                         TABLE IV                                       Run                A   B   C               PROCESS CONDITIONS                   Reactor Temperature, ° C. (° F.)   427 (800)    454 (850)    482 (900)        Catalyst-to-Oil Ratio   6.5   6.1   5.9       Hydrocarbon Partial Pressure,   117 (17.0)   114 (16.5)   122 (17.7)       kPa (psia)       System Pressure, kPa (psig)   278 (40.3)   276 (40.0)   273 (39.6)       LHSV, hr −1     4.6   4.6   4.6       PRODUCT YIELDS, wt-%       Dry Gas   0.4   0.7   1.1       C 3 &#39;s   1.6   2.4   3.4       C 4 &#39;s   6.1   7.8   9.4       C 5   +  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                  
 
         [0053]    The foregoing qualities and yields pertaining to the C 5 +gasoline have been adjusted to reflect the fact that C 4 &#39;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 C 4   −  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.  
         [0054]    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.  
                                                               TABLE V                           Gasoline Composition       Full Range Feed vs. Product            Carbon #   Total   Naphthenes   Isoparaffins   n-Paraffins   Cyclic-Olefins   Iso-Olefins   n-Olefins   Aromatics                5   24.93 (25.32)    0.1 (0.0)    7.33 (17.48)   1.53 (2.49)   0.61 (0.14)    8.24 (3.17)   7.12 (2.03)   — (—)        6   23.00 (23.92)   1.22 (1.79)    6.31 (15.28)   0.92 (1.63)   2.04 (0.32)    7.33 (3.12)   4.68 (1.11)    0.51 (0.67)        7   18.17 (16.43)   1.79 (1.94)    3.87 (7.59)   0.51 (0.90)   2.24 (0.29)    4.88 (1.33)   2.54 (0.29)    2.34 (4.10)        8   14.96 (14.41)   1.53 (0.87)    2.54 (3.88)   0.51 (0.63)   1.02 (—)    3.15 (0.46)   1.32 (—)    4.88 (8.57)        9   12.72 (17.58)   0.92 (0.64)    1.83 (2.31)   0.41 (0.43)   0.31 (—)    1.83 (0.16)   0.81 (—)   A9 +       10    3.47 (1.79)   0.32 (0.24)    1.12 (1.10)   0.41 (0.46)   0.00 (—)    1.12 (—)   0.51 (—)    6.61 (14.04)       11    1.42 (0.52)   — (—)    0.51 (0.54)   0.31 (0.00)   0.00 (—)    0.41 (—)   0.2 (—)       Total   98.68 (100)   5.87 (5.49)   23.51 (48.19)   4.58 (6.55)   6.21 (0.74)   26.97 (8.23)   17.2 (3.42)   14.35 (27.37)            12   C 12   +  Non-Aromatics: 1.3                  
 
         [0055]    With regard to Table V, aromatics with nine or more carbon numbers are grouped together. Therefore, the numbers given for carbon numbers 10 and 11 in the “Total” column include only non-aromatic C 10 &#39;s and C 11 &#39;s. The minimal changes in total concentration of each carbon number fraction, especially in the C 5 -C 8  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.