Patent Publication Number: US-8124020-B2

Title: Apparatus for preventing metal catalyzed coking

Description:
FIELD OF THE INVENTION 
     This invention generally relates to an apparatus and process for producing desired products, such as light olefins including propylene. 
     DESCRIPTION OF THE RELATED ART 
     Fluid catalytic cracking (FCC) is a catalytic hydrocarbon conversion process accomplished by contacting heavier hydrocarbons in a fluidized reaction zone with a catalytic particulate material. The reaction in catalytic cracking, as opposed to hydrocracking, is carried out in the absence of substantial added hydrogen or the consumption of hydrogen. As the cracking reaction proceeds substantial amounts of highly carbonaceous material referred to as coke are deposited on the catalyst to provide coked or spent catalyst. Vaporous lighter products are separated from spent catalyst in a reactor vessel. Spent catalyst may be subjected to stripping over an inert gas such as steam to strip entrained hydrocarbonaceous gases from the spent catalyst. A high temperature regeneration with oxygen within a regeneration zone operation burns coke from the spent catalyst which may have been stripped. Various products may be produced from such a process, including a gasoline product and/or light product such as propylene and/or ethylene. 
     In such processes, a single reactor or a dual reactor can be utilized. Although additional capital costs may be incurred by using a dual reactor apparatus, one of the reactors can be operated to tailor conditions for maximizing products, such as light olefins including propylene and/or ethylene. 
     It can often be advantageous to maximize yield of a product in one of the reactors. Additionally, there may be a desire to maximize the production of a product from one reactor that can be recycled back to the other reactor to produce a desired product, such as propylene. 
     Much of the focus of FCC technology development over the past few years has been in maximizing propylene selectivity. This has driven most FCC technology licensors to develop a dual-riser FCC technology offering in which the primary feedstock, typically, VGO, is fed to one riser and a recycle stream of C 10 —, or any fraction thereof is recycled to a secondary riser. In this fashion, the primary riser and secondary riser can be operated in different modes to promote the most overall selective net yields. In typical operation, the primary riser would be operated less severely than the secondary riser. The secondary riser would be operated much more severely, to promote the formation of light olefins such as butylene, propylene and ethylene favored by higher temperature in the typical range of 538° to 593° C. (1000° to 1100° F.) and lower hydrocarbon partial pressure of less than 138 kPa (absolute) (20 psia). Feedstock to the secondary riser may be an FCC recycle or C 10 — material from other process units. 
     Those who have commercialized dual riser technology in the service of recycling naphtha to the secondary riser have all suffered from excessive coke formation in the secondary riser which has resulted in limited operating capability for these processes. In the known cases, operation was limited to weeks rather than months of operation before the unit had to be shut down and the coke removed. Thus, there is a need to provide a dual reactor apparatus for catalytic cracking that can avoid excessive coke formation in the secondary riser. 
     SUMMARY OF THE INVENTION 
     We have discovered that the excessive coking in the secondary reactor is due to Metal Catalyzed Coking (MCC). MCC is inhibited in conventional FCC units because sulfur species that decompose to form hydrogen sulfide in an FCC riser are sufficiently present in the hydrocarbon feed to an FCC unit. Hydrogen sulfide subsequently passivates the active metals in the FCC unit. We propose a process and apparatus of adding a sulfiding agent to an FCC riser or other reactor when hydrogen sulfide is insufficiently present to inhibit MCC. The sulfur species in the sulfiding agent is provided as hydrogen sulfide or provides a source of hydrogen sulfide, either by decomposition, liberation, or other chemical reaction, that subsequently forms a metal sulfide layer on the interior metal surface of the reactor internals. The layer of metal sulfide isolates the vapor phase coke precursors from the active metal sites on the internal surface to inhibit coking. 
    
    
     
       BRIEF DESCRIPTION OF THE DRAWING 
       The FIGURE is a schematic drawing of the present invention. 
     
    
    
     DETAILED DESCRIPTION OF THE DRAWING 
     MCC is characterized by a deposition of carbonaceous solids on hot metal surfaces and develops in processes in excess of 400° C., with a peak filamentous carbon formation rate in the range of about 550° to about 600° C. MCC can be a function of thermal decomposition, or catalytic reaction with the active metals and can have a considerable impact on a number of commercial processes including, catalytic steam reforming of methane, steam cracking of paraffinic feed stocks and processes involving carbon monoxide disproportionation reactions. It is well known that certain metals can increase the overall MCC deposition rate by catalyzing the growth of filamentous and graphitic types of deposits. The highest catalytic activity for carbon deposition is exhibited by iron, cobalt and nickel, and alloys containing these metals. An overall catalytic reaction pathway for MCC is generally believed to be the adsorption of ethylene, propylene or butylene onto a metal surface. The adsorbed light olefin then undergoes further dehydrogenation conversion to aromatics and alkyl aromatics which further condense until coke is formed. 
     Typical FCC reactions operate in the range of about 500° to about 600° C., which corresponds to the peak reaction rate for filamentous carbon formation. The most active metals identified to promote MCC are present in an FCC unit. The active hydrocarbon species that promote filamentous carbon formation are ethylene, propylene and butylene which are the target products from high propylene producing FCC technologies. Consequently, we believe the coking problem in secondary FCC riser processes is attributed to MCC. 
     MCC has not historically been observed in FCC operations. Most FCC units process feed stocks with substantial quantities of sulfur, typically about 0.1 to about 1.0 wt-%. Sulfur present in FCC feed decomposes to hydrogen sulfide which adsorbs on the metal surface to form a metal sulfide layer which isolates gas phase coke precursors from active metal sites on internal FCC reactor surfaces, thereby mitigating coke formation. We have found that in recycle streams the hydrogen sulfide generated by cracking the primary FCC feed is not typically present in the naphtha feed recycled to secondary FCC riser. Organic sulfur in the primary FCC products distributes preferentially to hydrogen sulfide and coke in the reaction products, then distributes preferentially into the heavier products, with the least amount of sulfur remaining in the naphtha and liquefied petroleum gas (LPG). In secondary risers processing naphtha, the naphtha can be largely deficient of contaminant sulfur, resulting in insufficient sulfide layering on the metal in the secondary riser to prevent MCC. Even if sulfur is present in the naphtha, unless it is of a form that will thermally decompose to form hydrogen sulfide, it will not form a layer to passivate the active metals that contribute to MCC. 
     We propose to add a sulfiding agent to a catalytic reactor to prevent MCC from causing a chronic coke problem in the secondary reactor. The sulfiding agent can be hydrogen sulfide or an organic sulfur compound that decomposes to hydrogen sulfide in a catalytic conversion environment and particularly a fluid catalytic cracking environment. The hydrogen sulfide can be provided in dry gas fed to the secondary reactor prior to amine treating. Hydrogen sulfide may also be provided by adding a commercially available SO x  scavenging additive, such as a magnesium aluminum oxide having a spinel structure, to the circulating catalyst inventory. The additive adsorbs SO x  in the oxidizing environment of the regenerator and desorbs hydrogen sulfide in the reducing environment of the reactor riser. However, the technical capability of using a SO x  additive to provide sufficient hydrogen sulfide content in the second reactor is highly dependent on the sulfur content of the feedstock to the first reactor. Preferred organic sulfur sources include commercially available sulfiding agents such as methyl sulfides like dimethyl sulfide (DMS) or dimethyl disulfide (DMDS), mercaptans and polysulfides which have been conventionally used in industrial practice as sulfiding agents for hydroprocessing units and pyrolysis furnaces. These organic sulfur sulfiding agents degrade into hydrogen sulfide in a fluid catalytic cracking and other reaction environments. Sulfur containing oils in the FCC product such as LCO, HCO and CSO are not preferred sulfiding agents because they are not expected to effectively thermally decompose to generate the quantities of hydrogen sulfide required to passivate the active metals. However, under certain conditions, these heavy FCC products may be effective. Lighter FCC products such as naphtha and LPG may also be effective sulfiding agents under certain conditions if sulfide compounds are not removed therefrom. 
     The addition of hydrogen sulfide bearing dry gas is preferably added to a fluidizing gas distributor or as an atomizing dispersion media to feed distributors for a riser reactor. The organic sulfur sulfiding agents may be added to a fluidizing gas distributor or preferably to the feed system any point upstream of the feed distributors. The maximum sulfur rate is not limited, but is suitably in the range of about 20 to about 2000 wppm and preferably about 50 to about 500 wppm relative to the fluids present in the reactor. The sulfiding agent should be added on a continuous basis because coking onset is very fast, and the sulfide will adsorb and desorb from the active metals on a continuous basis. 
     The present invention may be described with reference to four components: a primary or first reactor  10 , a regenerator vessel  60 , a product fractionation section  90  and a second reactor  170 . Many configurations of the present invention are possible, but a specific embodiment is 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 if the first and second reactors  10 ,  170  are not FCC reactors, one or both of the regenerator vessel  60  and the product fractionation section  90  may be optional. Additionally, the invention may be embodied in a single FCC reactor  170 . 
     The FIGURE shows the first reactor  10  which may be an FCC reactor that includes a first reactor riser  12  and a first reactor vessel  20 . A regenerator catalyst pipe  14  in upstream communication with the first reactor riser  12  meaning that that material flow is permitted from the regenerator catalyst pipe  14  to the first reactor riser  12 . Communication means that material flow is permitted between enumerated regions. The regenerator catalyst pipe  14  delivers regenerated catalyst from the regenerator vessel  60  at a rate regulated by a control valve  16  to the reactor riser  12  through a regenerated catalyst inlet. A fluidization medium such as steam from a distributor  18  urges a stream of regenerated catalyst upwardly through the first reactor riser  12  at a relatively high density. A plurality of feed distributors  22  in upstream communication with the first reactor riser  12  inject a first hydrocarbon feed  8 , preferably with an inert atomizing gas such as steam, across the flowing stream of catalyst particles to distribute hydrocarbon feed to the first reactor riser  12 . Upon contacting the hydrocarbon feed with catalyst in the first reactor riser  12  the heavier hydrocarbon feed cracks to produce lighter gaseous first cracked products while conversion coke and contaminant coke precursors are deposited on the catalyst particles to produce coked catalyst. 
     A conventional FCC feedstock and higher boiling hydrocarbon feedstock are a suitable first feed  8  to the first FCC reactor. The most common of such conventional feedstocks is a “vacuum gas oil” (VGO), which is typically a hydrocarbon material having a boiling range of from 343° to 552° C. (650° to 1025° F.) prepared by vacuum fractionation of atmospheric residue. Such a fraction is generally low in coke precursors and heavy metal contamination which can serve to contaminate catalyst. Heavy hydrocarbon feedstocks to which this invention may be applied include heavy bottoms from crude oil, heavy bitumen crude oil, shale oil, tar sand extract, deasphalted residue, products from coal liquefaction, atmospheric and vacuum reduced crudes. Heavy feedstocks for this invention also include mixtures of the above hydrocarbons and the foregoing list is not comprehensive. Usually, the first feed  8  has a temperature of about 140 to about 320° C. Moreover, additional amounts of feed may also be introduced downstream of the initial feed point. 
     The first reactor vessel  20  is in downstream communication with the first reactor riser  12  meaning that material flow is permitted from the first reactor riser  12  to the first reactor vessel  20 . The resulting mixture of gaseous product hydrocarbons and spent catalyst continues upwardly through the first reactor riser  12  and are received in the first reactor vessel  20  in which the spent catalyst and gaseous product are separated. A pair of disengaging arms  24  may tangentially and horizontally discharge the mixture of gas and catalyst from a top of the first reactor riser  12  through one or more outlet ports  26  (only one is shown) into a disengaging vessel  28  that effects partial separation of gases from the catalyst. A transport conduit  30  carries the hydrocarbon vapors, including stripped hydrocarbons, stripping media and entrained catalyst to one or more cyclones  32  in the first reactor vessel  20  which separates spent catalyst from the hydrocarbon gaseous product stream. The disengaging vessel  28  is partially disposed in the first reactor vessel  20  and can be considered part of the first reactor vessel  20 . Gas conduits  34  deliver separated hydrocarbon gaseous streams from the cyclones  32  to a collection plenum  36  in the first reactor vessel  20  for passage to a product line  88  via an outlet nozzle  38  and eventually into the product fractionation section  90  for product recovery. Diplegs  40  discharge catalyst from the cyclones  32  into a lower bed  42  in the first reactor vessel  20 . The catalyst with adsorbed or entrained hydrocarbons may eventually pass from the lower bed  42  into an optional stripping section  44  across ports  46  defined in a wall of the disengaging vessel  28 . Catalyst separated in the disengaging vessel  28  may pass directly into the optional stripping section  44  via a bed  48 . A fluidizing distributor  50  delivers inert fluidizing gas, typically steam, to the stripping section  44 . The stripping section  44  contains baffles  52  or other equipment to promote contacting between a stripping gas and the catalyst. The stripped spent catalyst leaves the stripping section  44  of the disengaging vessel  28  of the first reactor vessel  20  with a lower concentration of entrained or adsorbed hydrocarbons than it had when it entered or if it had not been subjected to stripping. The spent catalyst, preferably stripped, leaves the disengaging vessel  28  of the first reactor vessel  20  through a spent catalyst conduit  54  and passes into the regenerator vessel  60  at a rate regulated by a slide valve  56 . 
     The first reactor riser  12  can operate at any suitable temperature, and typically operates at a temperature of about 150° to about 580° C., preferably about 520° to about 580° C. at the riser outlet  24 . In one exemplary embodiment, a higher riser temperature may be desired, such as no less than about 565° C. at the riser outlet port  24  and a pressure of from about 69 to about 517 kPa (gauge) (10 to 75 psig) but typically less than about 275 kPa (gauge) (40 psig). The catalyst-to-oil ratio, based on the weight of catalyst and feed hydrocarbons entering the bottom of the riser, may range up to 30:1 but is typically between about 4:1 and about 10:1 and may range between 7:1 and 25:1. Hydrogen is not normally added to the riser. Steam may be passed into the first reactor riser  12  and first reactor vessel  20  equivalent to about 2-35 wt-% of feed. Typically, however, the steam rate will be between about 2 and about 7 wt-% for maximum gasoline production and about 10 to about 15 wt-% for maximum light olefin production. The average residence time of catalyst in the riser may be less than about 5 seconds. 
     The catalyst in the first reactor  10  can be a single catalyst or a mixture of different catalysts. Usually, the catalyst includes two components or catalysts, namely a first component or catalyst, and a second component or catalyst. Such a catalyst mixture is disclosed in, e.g., U.S. Pat. No. 7,312,370 B2. Generally, the first component may include any of the well-known catalysts that are used in the art of FCC, such as an active amorphous clay-type catalyst and/or a high activity, crystalline molecular sieve. Zeolites may be used as molecular sieves in FCC processes. Preferably, the first component includes a large pore zeolite, such as a Y-type zeolite, an active alumina material, a binder material, including either silica or alumina, and an inert filler such as kaolin. 
     Typically, the zeolitic molecular sieves appropriate for the first component have a large average pore size. Usually, molecular sieves with a large pore size have pores with openings of greater than about 0.7 nm in effective diameter defined by greater than about 10, and typically about 12, member rings. Pore Size Indices of large pores can be above about 31. Suitable large pore zeolite components may include synthetic zeolites such as X and Y zeolites, mordenite and faujasite. A portion of the first component, such as the zeolite, can have any suitable amount of a rare earth metal or rare earth metal oxide. 
     The second component may include a medium or smaller pore zeolite catalyst, such as a MFI zeolite, as exemplified by at least one of ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38, ZSM-48, and other similar materials. Other suitable medium or smaller pore zeolites include ferrierite, and erionite. Preferably, the second component has the medium or smaller pore zeolite dispersed on a matrix including a binder material such as silica or alumina and an inert filler material such as kaolin. The second component may also include some other active material such as Beta zeolite. These compositions may have a crystalline zeolite content of about 10 to about 50 wt-% or more, and a matrix material content of about 50 to about 90 wt-%. Components containing about 40 wt-% crystalline zeolite material are preferred, and those with greater crystalline zeolite content may be used. Generally, medium and smaller pore zeolites are characterized by having an effective pore opening diameter of less than or equal to about 0.7 nm, rings of about 10 or fewer members, and a Pore Size Index of less than about 31. Preferably, the second catalyst component is an MFI zeolite having a silicon to aluminum ratio greater than about 15, preferably greater than about 75. In one exemplary embodiment, the silicon to aluminum ratio can be about 15:1 to about 35:1. 
     The total mixture in the first reactor  10  may contain about 1 to about 25 wt-% of the second component, namely a medium to small pore crystalline zeolite with greater than or equal to about 1.75 wt-% of the second component being preferred. When the second component contains about 40 wt-% crystalline zeolite with the balance being a binder material, an inert filler, such as kaolin, and optionally an active alumina component, the mixture may contain about 4 to about 40 wt-% of the second catalyst with a preferred content of at least about 7 wt-%. The first component may comprise the balance of the catalyst composition. In some preferred embodiments, the relative proportions of the first and second components in the mixture may not substantially vary throughout the first reactor  10 . The high concentration of the medium or smaller pore zeolite as the second component of the catalyst mixture can improve selectivity to light olefins. In one exemplary embodiment, the second component can be a ZSM-5 zeolite and the mixture can include about 4 to about 10 wt-% ZSM-5 zeolite excluding any other components, such as binder and/or filler. 
     The regenerator vessel  60  is in downstream communication with the first reactor vessel  20 . In the regenerator vessel  60 , coke is combusted from the portion of spent catalyst delivered to the regenerator vessel  60  by contact with an oxygen-containing gas such as air to provide regenerated catalyst. The regenerator vessel  60  may be a combustor type of regenerator as shown in the FIGURE, which may use hybrid turbulent bed-fast fluidized conditions in a high-efficiency regenerator vessel  60  for completely regenerating spent catalyst. However, other regenerator vessels and other flow conditions may be suitable for the present invention. The spent catalyst conduit  54  feeds spent catalyst to a first or lower chamber  62  defined by an outer wall through a spent catalyst inlet. The spent catalyst from the first reactor vessel  20  usually contains carbon in an amount of from 0.2 to 2 wt-%, which is present in the form of coke. Although coke is primarily composed of carbon, it may contain from 3 to 12 wt-% hydrogen as well as sulfur and other materials. An oxygen-containing combustion gas, typically air, enters the lower chamber  62  of the regenerator vessel  60  through a conduit and is distributed by a distributor  64 . As the combustion gas enters the lower chamber  62 , it contacts spent catalyst entering from spent catalyst conduit  54  and lifts the catalyst at a superficial velocity of combustion gas in the lower chamber  62  of perhaps at least 1.1 m/s (3.5 ft/s) under fast fluidized flow conditions. In an embodiment, the lower chamber  62  may have a catalyst density of from 48 to 320 kg/m 3  (3 to 20 lb/ft 3 ) and a superficial gas velocity of 1.1 to 2.2 m/s (3.5 to 7 ft/s). The oxygen in the combustion gas contacts the spent catalyst and combusts carbonaceous deposits from the catalyst to at least partially regenerate the catalyst and generate flue gas. 
     The mixture of catalyst and combustion gas in the lower chamber  62  ascend through a frustoconical transition section  66  to the transport, riser section  68  of the lower chamber  62 . The riser section  68  defines a tube which is preferably cylindrical and extends preferably upwardly from the lower chamber  62 . The mixture of catalyst and gas travels at a higher superficial gas velocity than in the lower chamber  62 . The increased gas velocity is due to the reduced cross-sectional area of the riser section  68  relative to the cross-sectional area of the lower chamber  62  below the transition section  66 . Hence, the superficial gas velocity may usually exceed about 2.2 m/s (7 ft/s). The riser section  68  may have a catalyst density of less than about 80 kg/m 3  (5 lb/ft 3 ). 
     The regenerator vessel  60  also includes an upper or second chamber  70 . The mixture of catalyst particles and flue gas is discharged from an upper portion of the riser section  68  into the Lipper chamber  70 . Substantially completely regenerated catalyst may exit the top of the transport, riser section  68 , but arrangements in which partially regenerated catalyst exits from the lower chamber  62  are also contemplated. Discharge is effected through a disengaging device  72  that separates a majority of the regenerated catalyst from the flue gas. In an embodiment, catalyst and gas flowing up the riser section  68  impact a top elliptical cap of the riser section  68  and reverse flow. The catalyst and gas then exit through downwardly directed discharge outlets of disengaging device  72 . The sudden loss of momentum and downward flow reversal cause a majority of the heavier catalyst to fall to the dense catalyst bed  74  and the lighter flue gas and a minor portion of the catalyst still entrained therein to ascend upwardly in the upper chamber  70 . Cyclones  75 ,  76  further separate catalyst from ascending gas and deposits catalyst through diplegs  77 ,  78  into dense catalyst bed  74 . Flue gas exits the cyclones  75 ,  76  through a gas conduit and collects in a plenum  82  for passage to an outlet nozzle  84  of regenerator vessel  60  and perhaps into a flue gas or power recovery system (not shown). Catalyst densities in the dense catalyst bed  74  are typically kept within a range of from about 640 to about 960 kg/m 3  (40 to 60 lb/ft 3 ). A fluidizing conduit delivers fluidizing gas, typically air, to the dense catalyst bed  74  through a fluidizing distributor  86 . In an embodiment, to accelerate combustion of the coke in the lower chamber  62 , hot regenerated catalyst from a dense catalyst bed  74  in the upper chamber  70  may be recirculated into the lower chamber  62  via recycle conduit  80 . 
     The regenerator vessel  60  may typically require 14 kg of air per kg of coke removed to obtain complete regeneration. When more catalyst is regenerated, greater amounts of feed may be processed in the first reactor  10 . The regenerator vessel  60  typically has a temperature of about 594° to about 704° C. (100° to 1300° F.) in the lower chamber  62  and about 649° to about 760° C. (1200° to 1400° F.) in the upper chamber  70 . The regenerated catalyst pipe  14  is in downstream communication with the regenerator vessel  60 . Regenerated catalyst from dense catalyst bed  74  is transported through regenerated catalyst pipe  14  from the regenerator vessel  60  back to the first reactor riser  12  through the control valve  16  where it again contacts feed as the FCC process continues. 
     In addition, the first reactor  10  can be operated at low hydrocarbon partial pressure in one desired embodiment. Generally, a low hydrocarbon partial pressure can facilitate the production of light olefins. Accordingly, the pressure in the first reactor riser  12  can be about 170 to about 250 kPa with a hydrocarbon partial pressure of about 35 to about 180 kPa, preferably about 70 to about 140 kPa. A relatively low partial pressure for hydrocarbon may be achieved by using steam as a diluent, in the amount of about 10 to about 55 wt-%, preferably to about 15 wt-% of the feed. Other diluents, such as dry gas, can be used to reach equivalent hydrocarbon partial pressures. 
     The first cracked products in the line  88  from the first reactor  10 , relatively free of catalyst particles and including the stripping fluid, exits the first reactor vessel  20  through the outlet nozzle  38 . The first cracked products stream in the line  88  may be subjected to additional treatment to remove fine catalyst particles or to further prepare the stream prior to fractionation. The line  88  transfers the first cracked products stream to the product fractionation section  90  that in an embodiment may include a main column  100  and a gas concentration section  114 . 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 in an overhead line  102 . The overhead stream in overhead line  102  is condensed in a condenser  104  and cooled in a cooler  106  before it enters a receiver  108 . A line  110  withdraws a light off-gas stream from the receiver  108 . The off-gas contains LPG and dry gas. The dry gas contains hydrogen sulfide which can serve as a sulfiding agent. A bottom liquid stream of light gasoline leaves the receiver  108  via a line  112 . Both lines  110  and  112  may be fed to the gas concentration section  114 . In the gas concentration section  114  many streams are separated such as by fractionation to generate a light olefins line  116 , a light naphtha line  118  and a dry gas line  120 . The dry gas stream may be concentrated predominantly into a hydrogen sulfide stream or may be part of a more comprehensive stream, but will be represented by dry gas line  120 . At least a portion of the dry gas stream is taken by recycle dry gas sulfiding agent line  122  to feed dry gas mixing sulfiding agent line  124  and/or dedicated dry gas sulfiding agent line  184 . The main column  100  also provides a heavy naphtha stream, a light cycle oil (LCO) stream and a heavy cycle oil (HCO) stream through lines  126 ,  128  and  130 , respectively. Parts of the streams in the lines  126 ,  128  and  130  are all circulated through heat exchangers  132 ,  134  and  136  and reflux loops  138 ,  140  and  142 , respectively, to remove heat from the main column  100 . Streams of heavy naphtha, LCO and HCO are transported from the main column  100  through respective lines  144 ,  146  and  148 . A clarified oil (CO) fraction may be recovered from the bottom of the main column  100  via a line  150 . Part of the CO fraction is recycled through a reboiler  152  and returned to the main column  100  through a line  154 . The CO stream is removed from the main column  100  via a line  156 . 
     The light naphtha fraction preferably has an initial boiling point (IBP) below about 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 about 127° C. (260° F.). The boiling points for these fractions are determined using the procedure known as ASTM D86-82. A portion of the light naphtha stream in light naphtha line  118  may be recovered in line  156  for further processing or storage and another portion in feed line  158  regulated by a control valve may be delivered to recycle feed line  166  for recycle as feed to the second reactor  170 . The heavy naphtha fraction has an IBP at or above about 127° C. (260° F.) and an EP at a temperature above about 200° C. (392° F.), preferably between about 204° and about 221° C. (400° and 430° F.), particularly at about 216° C. (420° F.). A portion of the heavy naphtha stream in line  144  may be recovered in line  160  for further processing or storage and another portion in line  162  regulated by a control valve may be delivered to recycle feed line  166  for recycle as feed to the second reactor  170 . The LCO stream has an IBP at about the EP temperature of the heavy naphtha and an EP in a range of about 260° to about 371° C. (500° to 700° F.) and preferably about 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 about 371° to about 427° C. (700° 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. 
     It is also contemplated that in the product recovery section  90  that a less refined separation of dry gas from LPG and/or naphtha streams may be performed to allow hydrogen sulfide containing dry gas to be added to the second reactor  170  in a hydrocarbon feed line containing the LPG and/or naphtha stream instead of by transport through a separate sulfiding agent line. 
     The second reactor  170  may be a second FCC reactor. Although the second reactor  170  is depicted as a second FCC reactor, it should be understood that any suitable reactor can be utilized, such as a fixed bed or a fluidized bed. The second hydrocarbon feed may be fed to the secondary FCC reactor in recycle feed line  166  via feed distributor line  168  and/or fluidizing feed line  172  and fluidizing distributor supply line  174 . The second feed can at least partially be comprised of C 10 — hydrocarbons and preferably C 4  to C 10  olefins. Preferably, the second hydrocarbon feed predominantly comprises hydrocarbons with 10 or fewer carbon atoms. Predominantly means over 50 wt-% and preferably over 80 wt-%. The second feed may comprise any hydrocarbon containing feed that is low in sulfur compounds that decompose to hydrogen sulfide such as a pyrolysis oil from a pyrolysis reactor, Fischer-Tropsch wax from a Fischer-Tropsch reactor, reformate from a catalytic reforming reactor, straight run naphtha from a crude column and animal fat and vegetable oils from an appropriate reactor or source. The second feed is preferably a portion of the first cracked products produced in the first reactor  10 , fractionated in the main column  100  of the product fractionation section  90  via recycle feed line  166  and provided to the second reactor  170 . In an embodiment, the second reactor is in downstream communication with the product fractionation section  90  and/or the first reactor  10  which is in upstream communication with the product fractionation section  90 . The second reactor  170  can include a second reactor riser  180 . The second hydrocarbon feed is contacted with catalyst delivered to the second reactor  170  by a catalyst return pipe  176  in upstream communication with the second reactor riser  180  to produce cracked upgraded products. 
     The present invention contemplates adding a sulfiding agent to the second reactor  170  to inhibit metal catalyzed coking therein. The recycle dry gas sulfiding agent line  122  is a dedicated source of a sulfiding agent in upstream communication with the second reactor riser  180 . In other words, dry gas and hydrogen sulfide would not be fed to the second reactor  170  except to prevent metal catalyzed coking because they will not convert to desirable hydrocarbon products and will have to be removed from the upgraded products exiting the second reactor  170 . The introduction of hydrocarbon feed and sulfiding agent to the second reactor  170  can be performed in several embodiments shown in the FIGURE. 
     In a first embodiment, the second hydrocarbon feed can be injected into a second reactor riser  180  by a feed distributor  178  in upstream communication with the second reactor riser  180  and in downstream communication with a feed distributor line  168  which is in downstream communication with recycle feed line  166 . Feed distributor line  168  may take a portion or all of the recycle feed stream from recycle feed line  166 . The recycle feed line  166  is in downstream communication with the overhead line  102  of the main column  100  which is in downstream communication with the first reactor  10 . The feed rate in feed distributor line  168  may be regulated by a control valve. The feed distributor  178  may be located above a fluidizing distributor  182  which is in upstream communication with the second reactor riser  180 . The fluidizing distributor  182  provides a fluidizing gas, such as steam and/or a light hydrocarbon, to the second reactor riser  180  to fluidize the catalyst. In such an embodiment, dry gas from recycle dry gas sulfiding agent line  122  may be independently added to the fluidizing distributor  182  in a base of the second reactor riser  180  via dedicated dry gas sulfiding agent line  184  in downstream communication with the recycle dry gas sulfiding agent line  122  and bypassing atomizing dry gas sulfiding agent line  186  in fluidizing sulfiding agent line  188  and fluidizing distributor supply line  174 . The dry gas thus serves both as a fluidizing gas and as a sulfiding agent added to the second reactor riser  180  of the second reactor  170 . The recycle dry gas sulfiding agent line  122 , the dedicated dry gas sulfiding agent line  184  and the fluidizing sulfiding agent line  188  are dedicated sources of a sulfiding agent in upstream communication with the fluidizing distributor  182  and the second reactor  170 . Dry gas bearing hydrogen sulfide in recycle dry gas sulfiding agent line  122 , dedicated dry gas sulfiding agent line  184  and fluidizing sulfiding agent line  188  can also be used as an inert fluidizing gas for other parts of the second reactor  170 . In this embodiment, control valves in feed lines  158  and/or  162  and  168  and in sulfiding agent lines  122 ,  184  and  188  may be open and control valves in feed lines  172  and sulfiding agent lines  124  and  186  may be closed. 
     In a second embodiment, when the second feed is liquid, a dry gas containing hydrogen sulfide may be added to the liquid second feed in the feed distributor  178  to atomize the liquid hydrocarbon second feed and passivate metals in the second reactor. The recycle dry gas sulfiding agent line  122  is a dedicated source of a sulfiding agent in upstream communication with the feed distributor  178  via atomizing dry gas sulfiding agent line  186 . Atomizing dry gas sulfiding agent line  186  in downstream communication with dedicated dry gas sulfiding agent line  184  provides dry gas to a gas inlet of the feed distributor  178 . Sulfiding agent may be added to the second reactor according to this embodiment in addition to or instead of the way sulfiding agent is added in the first embodiment; i.e., by addition through the fluidizing distributor  182 . Consequently, opening of control valve in line  186  in addition to the control valves opened and closed in other embodiments will allow operation according to this second embodiment. Accordingly, at least the control valves in sulfiding agent lines  122 ,  184  and  186  must be opened to operate under this embodiment. 
     In a third embodiment, essentially all of the second hydrocarbon feed in recycle feed line  166 , i.e., at least about 90%, by mole is in a gas phase. Generally, the temperature of the second hydrocarbon feed can be about 120° to about 600° C. when entering the second reactor riser  180  and, preferably, at least be above the boiling point of the components. In this embodiment, the second hydrocarbon feed can be fed directly to the fluidizing distributor  182  in the base of the second riser to fluidize the catalyst and to feed the second reactor riser  180 . In this embodiment, shown in the FIGURE, one or all of control valves in sulfiding agent lines  122  and  124  and feed lines  158  and/or  162  and  172  are open to allow dry gas containing hydrogen sulfide in recycle dry gas sulfiding agent line  122  and dry gas mixing sulfiding agent line  124  and light naphtha in light naphtha line  158  and/or heavy naphtha in heavy naphtha line  162  to recycle as secondary feed in recycle feed line  166 , fluidizing feed line  172  and fluidizing distributor supply line  174  to be distributed to the riser by fluidizing distributor  182 . Valves in feed line  168  and sulfiding agent lines  184 ,  186  and  188  may typically be closed in this embodiment. The dry gas should contain sufficient hydrogen sulfide to passivate the metals that can catalyze coking in the second reactor riser  180  of the second reactor  170 . A heat exchanger  190  may be necessary on fluidizing feed line  172  to vaporize the recycled secondary feed. In this embodiment, fluidizing distributor supply line  174  serves as a feed line and the fluidizing distributor  182  serves as a feed distributor. 
     Hydrogen sulfide, in dry gas or not, or organic sulfur additives such as methyl sulfides, mercaptans and polysulfides may be suitable additive sulfiding agents that are added to the second reactor  170 . The additive sulfiding agents may be added to the second feed in feed lines  158 ,  162 ,  166 ,  168 ,  172  or  174  or elsewhere upstream of the second reactor  170 . For example, additive sulfiding agent line  192  may add a sulfiding agent directly to the fluidizing feed line  172 . Sulfiding agents may also be added directly to the second reactor riser  180 , to fluidizing gas upstream of the fluidizing distributor  182  or even to the catalyst entering the riser in catalyst return pipe  176 . If a SO x  scavenger additive is added to the catalyst, hydrogen sulfide adsorbed on the additive may be delivered to the second reactor  170  via pipe  216  and catalyst return pipe  176 , making one or both of the catalyst return pipe  176  and pipe  216  a sulfiding agent line. The sulfiding agent stream in the sulfiding agent line preferably has a concentration of at least 1000 wppm of hydrogen sulfide or a compound that can convert to hydrogen sulfide in the reactor environment. The concentration of sulfur relative to the fluids in the second reactor  170  should be maintained to be at least about 20 wppm and preferably about 50 wppm. In a riser reactor, the concentration of sulfur should be maintained to be at least about 20 wppm and preferably about 50 wppm relative to the hydrocarbon and inert gases in the reactor. In an embodiment, the concentration of sulfur relative to the fluids in the second reactor should be maintained to be no more than about 2000 wppm and preferably no more than about 500 wppm. In a riser reactor, the concentration of sulfur should be maintained to be no more than about 2000 wppm and preferably no more than about 500 wppm relative to the hydrocarbon and inert gases in the reactor. 
     The sulfiding agent lines  122 ,  124 ,  176 ,  184 ,  186 ,  188  and  192  are distinct from the feed lines  158  and  162 . When the control valve in line  124  is closed, lines  166 ,  168  and  172  are also feed lines from which sulfiding agent lines  122 ,  184 ,  186  and  188  are distinct. When control valves in lines  124  and  172  are closed, fluidizing feed line  172  no longer carries feed but fluidizing distributor supply line  174  becomes a sulfiding agent line from which feed lines  158 ,  162 ,  166  and  168  are distinct. Although the streams in the sulfiding agent lines and feed lines may be mixed in a downstream location, these streams are separate from each other in at least an upstream location. Accordingly, sulfiding agent lines provide a sulfiding agent that is separate from the second hydrocarbon feed upstream of the second reactor  170 . 
     Generally, the second reactor  170  may operate under conditions to convert the hydrocarbon feed to smaller hydrocarbon products. C 10 — olefins crack into one or more light olefins, such as ethylene and/or propylene. A second reactor vessel  194  is in downstream communication with the second reactor riser  180  for receiving upgraded products and catalyst from the second reactor riser. The mixture of gaseous, upgraded product hydrocarbons and catalyst continues upwardly through the second reactor riser  180  and is received in the second reactor vessel  194  in which the catalyst and gaseous hydrocarbon, upgraded products are separated. A pair of disengaging arms  196  may tangentially and horizontally discharge the mixture of gas and catalyst from a top of the second reactor riser  180  through one or more outlet ports  198  (only one is shown) into the second reactor vessel  194  that effects partial separation of gases from the catalyst. The catalyst can drop to a dense catalyst bed  200  within the second reactor vessel  194 . Afterwards, the upgraded hydrocarbon products can be separated from the catalyst and be removed from the second reactor  170  through an outlet  204  in downstream communication with the second reactor  170  through an upgraded products line  206 . The upgraded products in upgraded products line  206  may be directed to one or more cyclones  32  in the first reactor vessel  20  of the first reactor  10 . These cyclones  32  may be dedicated just to the upgraded products from the second reactor  170  with a dedicated line (not shown) to the product fractionation section  90  or specifically the gas concentration section  114  or may just mix with the products from the first reactor riser  12  and travel together to the product fractionation section  90  in line  88 . Alternatively, the second reactor vessel  194  may contain or have one or more cyclones to further separate gaseous upgraded products from catalyst and travel via upgraded products line  206  to the gas concentration section  114  of the product fractionation section  90 . Upgraded products line  206  may alternatively deliver upgraded products to line  88  for transport to the main column  100  of the product fractionation section  90 . 
     In some embodiments, the second reactor  170  can contain a mixture of the first and second catalyst components as described above. In one preferred embodiment, the second reactor  170  can contain less than about 20 wt-%, preferably about 5 wt-% of the first component and at least 20 wt-% of the second component. In another preferred embodiment, the second reactor  170  can contain only the second component, preferably a ZSM-5 zeolite, as the catalyst. 
     Separated catalyst may be recycled via a recycle catalyst pipe  208  from the second reactor vessel  194  regulated by a control valve  210  back to the second reactor riser  180  to be contacted with the second feed. Optionally, catalyst can be provided from the stripping section  44  of the first FCC reactor via a pipe  214  and/or the regenerator vessel  60  via a pipe  216  both regulated by control valves to the second reactor  170 . Both pipes  214  and  216  may be in upstream communication with the recycle catalyst pipe  208 . Catalyst return pipe  176  may be a part of the recycle catalyst pipe  208 . In an embodiment, catalyst from the second reactor vessel  194  is delivered by pipe  202  to the first reactor, preferably to the stripping section  44 , and is delivered, preferably after stripping, via spent catalyst conduit  54  to the regenerator vessel  60  for regeneration. Regenerated catalyst may be returned by pipe  216  back to the base of the second reactor riser  180  via catalyst return pipe  176 . In this embodiment, the catalyst in the first and second reactors  10  and  170  are mixed and may be of uniform composition in both reactors. 
     In another embodiment, the second reactor  170  is isolated from the regenerator vessel  60 , so that regenerated catalyst is only returned to the first reactor  10  and the second reactor  170  does not send catalyst to the regenerator vessel  60  or receive regenerated catalyst therefrom. In this embodiment, the second catalyst component, by not being exposed to repeated regenerations, retains more of its activity. Instead, the second catalyst component can be added to the second reactor  170  and the catalyst in the second reactor vessel  194  can be periodically or continuously dispensed through the pipe  202  regulated by a control valve to the stripping section  44  of the first reactor  10 . The dispensed catalyst can combine with the catalyst in the first reactor  10  and provide additional catalyst activity therein. Fresh catalyst can replace dispensed catalyst to maintain activity in the second reactor  170 . 
     The second reactor riser  180  can operate in any suitable condition, such as a temperature of about 425° to about 705° C., preferably a temperature of about 550° to about 600° C., and a pressure of about 40 to about 700 kPa, preferably a pressure of about 40 to about 400 kPa, and optimally a pressure of about 200 to about 250 kPa. Typically, the residence time of the second reactor riser  180  can be less than about 5 seconds and preferably is between about 2 and about 3 seconds. Exemplary risers and/or operating conditions are disclosed in, e.g., US 2008/0035527 A1 and U.S. Pat. No. 7,261,807 B2. 
     Without further elaboration, it is believed that one skilled in the art can, using the preceding description, utilize the present invention to its fullest extent. The preceding preferred specific embodiments are, therefore, to be construed as merely illustrative, and not limitative of the remainder of the disclosure in any way whatsoever. 
     In the foregoing, all temperatures are set forth in degrees Celsius and, all parts and percentages are by weight, unless otherwise indicated. 
     From the foregoing description, one skilled in the art can easily ascertain the essential characteristics of this invention and, without departing from the spirit and scope thereof, can make various changes and modifications of the invention to adapt it to various usages and conditions.