US 3849291 A
Description (OCR text may contain errors)
United States Patent [191- Owen [ Nov, 19, 1974 HIGH TEMPERATURE CATALYTIC CRACKING WITH LOW COKE PRODUCING CRYSTALLINE ZEOLITE CATALYSTS  Inventor: Hartley Owen, Belle Mead, NJ.
 Assignee: Mobil Oil Corporation, New York,
22 Filed: Oct.5, 1971 21 App1.No.: 186,639
3,679,576 7/1972 McDonald 208/74 3,702,886 11/1972 Argauer et a1 423/328 3,706,654 12/1972 Bryson et a1. 208/74 3,748,251 7/1973 Demmel et a1. 208/74 3,758,403 9/1973 Rosinski et al. 208/120 3,764,520 10/1973 Kimberlin et a]. 208/1 11 3,769,202 10/1973 Plank et a1 208/111 OTHER PUBLICATIONS Voorhies Carbon Formation in Catalytic Cracking," Ind. Eng. Chem. 37, 318-322, (1945). v Shankland and Schmitkons, Determination of Activity and Selectivity of Cracking Catalyst." Proc. AP] 27, (111), 1947, pp. 57-77.
Primary Examiner-Delbert E. Gantz Assistant Examiner-G. E. Schmitkons Attorney, Agent, or Firm-Andrew L. Gaboriault; Carl D. Famsworth s7 ABSTRACT A method and arrangement of catalyst handling steps is described for practicing a selective high temperature catalytic cracking of hydrocarbons which will take advantage of the low coke producing catalyst of high activity and selectivity.
16 Claims, 5 Drawing Figures  References Cited UNITED STATES PATENTS 2,461,958 2/1949 Bonnell 208/74 2,688,401 9/1954 Schmitkons et a1. 209/172 2,847,364 8/1958 Hirsch 208/164 2,883,332 4/1969 Wickham 208/74 2,892,773 6/1959 Hirsch et a1 208/213 2,900,324 8/1969 Patton et al. 208/74 3,494,858 2 1970 Lyckenbach ..'..208/164 3,661,800 5/1972 Pfeiffer eta 1. 252/417 50 Reg i Air J v x 2 r 6 J Q D iluenr Coke on charge, 7 w? .JSVISISH PATENTEL g 1 0r 3 Figure l Coke vs.Conversi0n After Thermal Treatment (Colcinecl 3hr/|200F/oir) Bench FCCTesTs: 9|5F, 5C/O, 5WHSV,WCMCGO Catalysts E, H, and O bench tested 01 2C/O, |2.5 WHSV mwwk CD 00000 Conversion, vol
Coke on charge, wt
Pmm 22v 1 9 I914 3.849.291
SHEET 3 FigurelI Coke vs. Conversion After SteomTreotrnen'r (4hr/I400 F/ O psig) Bench FCC T esfs: 925F, 5 0/0, 5WHSV,WCMCGO Cofolysfs E, I and R bench tested at 3 0/0, 8.3 WHSV Conversion,"/ vol BACKGROUND OF THE INVENTION produce gasoline product. Thus considerable difficulty has been encountered in the prior systems in obtaining high yields of conversion products without excess production of the carbonaceous contaminants.
The present trend in catalytic cracking operations is concerned with those systems which will use more active and selective cracking catalysts such as those comprising crystalline zeolites for performing the conversion of one or more high boiling hydrocarbon fractions of the same or different boiling range and coke producing characteristics to gasoline boiling range products. Thus crystalline zeolite cracking technology necessarily requires using much more sophisticated cracking systems than those known or disclosed in the prior art in order to take fully advantage of the catalysts conversion capabilities. Many prior art systems and those converted for the use of high activity crystalline zeolite cracking catalysts have produced an inefficient operation causing undue catalyst regeneration, excessive re cycle of unconverted charge and generally inefficient use of the catalyst composition.
The invention defined herein is concerned with an improved sequence of conversion steps which will more efficiently utilize the capabilities of a crystalline zeolite cracking catalyst of high activity and high selectivity.
SUMMARY OF THE INVENTION This invention relates to the catalytic cracking of hydrocarbon oils having finely divided catalyst particles suspended in gasiform material comprising hydrocarbon reactant material. It relates more particularly to the catalytic cracking of selected hydrocarbon oils or fractions thereof with suspended catalyst particles of high activity and selectivity-under conditions selected to produce gasoline and improve upon the heat balance of the operation.
In a more particular aspect, the method and arrangement of processing steps of this invention is designed to make maximum use of low coke forming type crystalline zeolite cracking catalyst in short contact time riser reactors at elevated cracking temperatures.
The present invention is concerned with a method of operation which will circumvent the problems associated with using high activity crystalline zeolite cracking catalysts at high temperatures and low catalyst to oil ratios resulting in a used catalyst of relatively low carbon content.
The essence of the method and system herein described comprises a dual riser fluid catalytic cracking system effected with freshly regenerated catalyst in a first riserreactor and a mixture of freshlyregenerated catalyst with catalyst separated from the conversion products of the first riser in a second riser reactor.
Thus, the present invention is particularly designed and directed to the use of a relatively high catalyst to oil ratio in a second riser reactor so as to increase the severity of the cracking operation therein over that employed in a first riser reactor in addition to providing more optimum utilization of the available crystalline zeolite cracking catalyst activity in the processing combination.
The method and arrangement of catalyst contact steps described herein may be used to advantage in the following respects:
a. A low coke producing high selective crystalline zeolite cracking catalyst may be used much more efficiently.
b. The hydrocarbon charge stock and regeneration gas preheat facilities may be reduced over that required in a prior art system using the same low coke producing catalyst. f
c. More efficient cracking of low molecular weight and higher molecular weight hydrocarbon fractions may be effected.
d. The processing of high molecular weight charge fractions and/or heavy recycle oils may be accomplished in a short contact time riser reactor at a higher or lower average temperature and a higher catalyst to oil ratio than used for processing a lower molecular weight oil charge in a separate riser reactor.
e. The high activity low coke producing catalyst may be more selectively employed for converting hydrocarbon charge materials at temperatures in the range of 900F. to 1200F. employing catalyst-oil residence time less than 15 seconds and catalyst to oil ratios as high as 25 to 1.
f. The conversion selectively of the catalyst is enhanced in the processing sequence by the use of diluent viding a major portion of the endothermic reaction heat requirements of the dual riser conversion system.
DESCRIPTION OF THE INVENTION The present invention relates to the catalytic conversion of hydrocarbons with a selective high activity CAS crystalline aluminosilicate catalytic composition under elevated temperature conversion conditions maintained within restricted catalyst-oil residence contact time and catalyst to oil ratios in order to produce gasoline. The present invention is concerned with and relates to the method and means for effecting a selective cracking of hydrocarbon charge materials to gasoline boiling product with a very selective low coke producing crystalline alumino-sili'cate catalyst composition. In the method and system of this invention the contact time between catalyst and hydrocarbon varies with the hydrocarbon charge passed to the selective cracking operation. Generally the cracking operation effected in a dispersed catalyst phase relation zone is restricted to orders of magnitude amounting to only a few seconds up to about 15 seconds and in most instances the contact time will be restricted depending upon composition of the hydrocarbon charge to within the range of 4 to 12 seconds. Thus the concepts essential to practicing the present invention includes the method and sequence of catalyst cascade which will permit employing cracking temperatures in the range of 900F. to about l,200F. at a number of different catalyst to oil ratios and contact times herein identified. Further salient features of the present concept include the use of low coke producing catalysts in the riser reactors, desired catalyst-oil suspension relationships in a relatively low catalyst inventory system, and maximizing the use of heat available in the system to effect the catalytic conversion desired.
The utilization of highly selective low coke providing catalyst compositions comprising selected crystalline aluminosilicate catalyst compositions particularly suitable for accomplishing the processing concept are herein discussed. The processing concepts of this invention include a restricted contact time between a suspension of high activity catalyst and hydrocarbon feed being converted before discharge of the suspension into suitable separation equipment. Separation equipment particularly suitable for this purpose comprises one or more cyclone separators'at the discharge end of transfer line reactors which will minimize the time for separating catalyst particles and hydrocarbon material without substantially cooling upon discharge from the transfer line cracking zone. The use of several small cyclones in series reduces the separation time even though such small cyclones are subject to the constraint of some oil-catalyst mixing.
Effecting the dispersed catalyst phase cracking operation at conditions of high temperature in the range of 900 to l,200F. and a short contact time less than 15 seconds as advocated herein provides a low catalyst inventory operation which is economically attractive and favors the production of superior product quality in improved quantity.
To enhance the product quality and product distribution or selectively encountered in the operation discussed herein, a relatively high boiling virgin feed hydrocarbon material is passed to a first riser reactor with freshly regenerated catalyst for a limited residence time therein. A relatively high coke making hydrocarbon material such as coker gas oil and recycle hydrocarbons of the cracking operation is passed to a second riser reaction in contact with a mixture of catalyst comprising catalyst used in the first reactor and freshly regenerated catalyst. A diluent fluid such as wet or dry recycle gas or stream may be used in either or both reactors to reduce hydrocarbon partial pressure and/or control contact time by increasing total velocity of vapors and coke particles in the reactor. The improved method of operation herein defined for converting low and high boiling hydrocarbons either in a vapor, liquid and/or a partially vaporized condition depending upon the boiling range of the hydrocarbon charge relies upon a gasifo rm d ilue nt material being mixed with the charge to control the hydrocarbon partial pressure. The diluent also assists in breaking up the hydrocarbon feed into relatively fine droplets which more uniformly dis tribute themselves in intimate contact with the fine catalyst particles. Uasiform diluent materials which may be employed in the processing combination of this invention with varying degrees of success include steam, light gaseous hydrcarbons known as dry gas (C and lighter hydrocarbons) or wet gaseous hydrocarbon streams such as those comprising C and C hydrocarbons.
A further important aspect going to the very essence of this invention concerns itself with an apparatus and catalyst system of greatly improved flexibility going beyond that heretofore provided for utilizing the highly selective catalyst and heat available therefrom in a manner which permits conversion of hydrocarbon charge material of different coke producing characteristics to more acceptable products with greater efficiency. Thus it is clearly evident from that expressed above that operating flexibility to achieve desired levels of conversion on a per pass basis is contemplated and such variation may be had by providing one or more, for example, a plurality of suitably spaced apart hydrocarbon feed inlet nozzles along the length of riser reactor processing fresh or virgin feed material. In this embodiment preheated diluent material is combined with the hot freshly regenerated catalyst to form a suitable high temperature suspension which is caused to flow through the riser or confined reaction zone to which point of inlet of the hydrocarbon feed to be converted. The hydrocarbon feed is mixed with the flowing high temperature suspension and converted during transverse of the remaining portion of the reaction zone to a suspension separator such as one or more cyclone separators. In any event, it is important that the suspension be separated substantially immediately after a pre-,
determined residence time so that over cracking will be avoided. Thus, the prior art methods of discharging the suspension into large disengaging vessels wherein separation is effected primarily by velocity reduction and hindered settling is to be avoided since such separating techniques encompass high temperature after cracking in a more dilute catalyst phase. While it is true that temperature quenching of the suspension may be effected after a predetermined residence time, such quenching operations are particularly to be avoided. Thus it is preferred that the suspension in the riser be discharged into highly efficient cyclone separator equipment so that separated catalyst of relatively high heat content may be recovered for use as herein provided.
In one arrangement of this invention the catalyst separated from the riser reactor processing fresh or virgin feed being of relatively high temperature approaching l,0O0F. is subjected to a limited stripping action without substantially reducing the temperature of the separated catalyst. The stripped catalyst containing a small amount of residual carbonaceous material because of the low coke producing characteristics of the catalyst ,at the high temperatures employed is then combined with a quantity of freshly regenerated catalyst in an amount sufficient to form a mixture of catalyst having an elevated temperature sufficient for converting recycle feed in a second riser reactor as herein provided. The suspension passed through the second riser reactor under cracking conversion conditions is discharged into suitable separator equipment arranged to effect quick separation of catalyst particles from hydrocarbon vapors. The hydrocarbon vapors of the first and second riser reaction zones are recovered and passed to suitable fl'lLliUllfillll equipment. The catalyst eparated from the second suspension is stripped and then regenerated in one or more regeneration steps which will be adequate for removing deposited carbonaceous material from the catalyst by burning thereby heating the catalyst to an elevated temperature suitable for effecting catalytic conversion reactions.
It is contemplated in the method and arrangement of processing steps of this invention of introducing the hydrocarbon reactant in a downstream portion of the riser reactor and relying upon gasiform material for transporting the cracking catalyst to the point of hydrocarbon reactant inlet. Thus a plurality of spaced apart hydrocarbon feed inlets is provided in each riser.
In yet another embodiment, the catalyst separated from the different riser reactors is collected as a dense fluid bed of catalyst in a vessel or hopper separated by a common separator baffle extending upwardly from the vessel bottom and provided with catalyst flow through slots or passageways so that by adjustment of catalyst bed level on either side of the baffle, catalyst may be caused to flow through the slots in either direction as desired to the catalyst bed exerting a lower pressure. This variation in the basic arrangement of catalyst flow provides a further flexiblity in the method of operating the system and such flexibility is most desirable in present day catalytic operations using the more active and selective conversion catalysts.
A still further embodiment of the present invention concerns itself with collecting the separated catalyst from all the riser reactors in a single dense fluid bed of catalyst wherein the catalyst particles are mixed and stripped of hydrocarbon vapors with a suitable stripping gas. In this particular embodiment, provisions are made for passing the stripped catalyst to one or more stages of regeneration. In addition, provisions are made for withstanding catalyst particles from either the upper or lower portion of this catalyst bed for admixture with fresh regenerated catalyst and use in the second riser reactor as discussed before.
It will be noted from an understanding of the above expressed concepts that in all of the embodiments contemplated and discussed herein, that one feature of the improved processing combination discussed herein is in the use of a high activity low coke producing catalyst composition and maximizing, with such a low coke producing catalyst, the recovery of heat available in the process by providing catalyst cascading provisions and techniques in the processing system. That is, the heat retained by the catalyst after the short time conversion of fresh feed is supplemented by admixture with freshly regenerated catalyst and thereafter used to process more refractory materials such as high molecular weight material fractions, coker gas oilsand recycle stocks.
In the method and processing system of this invention it is also contemplated employing catalyst compositions in admixture with one another as separate discrete particles or combined as a single particle of catalyst. Thus, the essential requirement of a low coke producing crystalline alumino silicate cracking catalyst composition may be combined with a metal oxide which will enhance regeneration of the catalyst by converting CO to CO thereby improving upon the available heat recovery from the system by the catalyst. Also the low coke producing cracking catalyst composition may be combined with a ZSM-S type of catalyst particle which is also of low coke producing characteristics as a homogeneous particle mixture or the different catalyst compositions may be used as separate and discrete particles in the system so that the amount of one charged to the system may be different than the amount of the other charged and circulated in the system. In one particular embodiment, it is contemplated combining the metal oxide such as chromium oxide with the ZSM-S type of catalyst particle since such a composition is more likely to contain less coke than the more conventional or low coke producing catalyst particles used therewith to ef fect primarily cracking of the hydrocarbon-charge to each riser reactor. Under the cracking temperature conditions employed particularly in the fresh feed riser reactor, olefinic constituents formed during the cracking step are cyclized to form aromatics particularly by the ZSM-S type catalyst. The ZSM-S type catalyst may be combined with the low coke producing zeolite catalyst in amounts ranging from about 10 up to about 90 percent but more usually will be in a range of from about 40 to about percent. The ZSM-S type of catalyst is more fully discussed hereinafter.
The method of operation herein defined operates to take advantage of heat available in regenerated catalyst, minimizes reaction time within desired limits and may be used to adjust the residual coke on catalyst eventually recycled to the regeneration zone. The amount of coke on the catalyst is controllable and controlled by using the catalyst separated from the products of the virgin feed conversion step to effect conversion of the recycle and coker gas oil in the second transfer line cracking step. Thus the catalyst used in the first transfer line reactor is cascaded to the second and the conversion heat requirements of the second is provided by feed preheat and by adding hot freshly regenerated catalyst to the catalyst cascaded thereto from the first or virgin feed conversion step.
The catalyst regeneration techniques envisioned for ,use in the concept of the present invention are those which will permit removal of hydrocarbonaceous material or residue carbonaceous material from the catalyst particles such as by burning in the presence of oxygen rich gas or air whereby the catalyst will be heated to an elevated temperature in the range of I,IOQF. up to about 1,400F. It is important to the operation of this invention that the catalyst be stable at the elevated temperatures above defined so as to provide a major v portion of the heat required in the catalytic cracking or conversion steps above discussed. Catalyst regeneration temperatures-above 1,200F. and up to l,400F. are therefore contemplated. The catalyst regeneration techniques envisioned for use with the cracking steps of this invention encompass riser or transfer lines regeneration zones either alone or in combination with dense fluidized catalyst bed regeneration techniques through which the catalyst moves as a suspension in oxygen containing regeneration gases under conditionsto maximize retention of generated heat in the catalyst. A combination of dispersed-catalyst-phase regeneration zones alone or in combination with a dense fluid catalyst bed operation may be employed,in which arrangement the catalyst and regeneration gases will move to achieve the desired removal of carbonaceous deposits. It is contemplated employing the sequence of riser dispersed catalyst phase operations through which the 'catalyst sequentially moves as a suspension in regeneration gas under regenerating temperature conditions with the amount of regeneration gas, such as air, being passed to each reactor being controlled or limited so as to limit temperatures encountered during regeneration of the catalyst. Regeneration temperatures may also be raised by combining a carbon monoxide conversion metal oxide with the catalyst as mentioned above and by increasing the amount of coke or hydrocarbonaceous material deposited on the catalyst during the cracking step before being passed to the regeneration step. Since air is a common regeneration gas, varying the amount of carbonaceous material on catalyst passed to the regenerator is most effective for temperature control. In another arrangement one or more riser regenerators may be used in conjunction with a more dense catalyst bed regeneration arrangement, such as provided by a dense fluid catalyst bed or the countercurrent flowof catalyst downwardly through a regeneration zone and counter-current to upwardly flowing regeneration gas. In such arrangements the more dense catalyst suspension as it moves downwardly through the system will encounter an ever increasing concentration of oxygen rich regeneration gas. In these systems and arrangement of regeneration zones, it is preferred for economic reasons that the equipment utilized be sized to minimize the catalyst inventory of the system consistent with providing a commercially attractive operating arrangement.
It is clear from the above discussion that a considerably greater utilization of available heat is provided in the cracking operation discussed and that more than one such combination of riser reactors may be employed with a common catalyst regeneration system. The arrangement herein discussed is particularly attractive for processing the same or a combination of hydrocarbon charge materials (i.e., heavy gas oil, light gas oil, recycle, straight run gasoline or naphtha) and each combination relies upon and maintains its own reaction parameters such as temperature, space velocity reaction time and catalyst-oil ratio. The arrangement is also attractive for multiple passing a hydrocarbon charge to be converted so as to increase gasoline product yields.
In the cracking operation discussed herein the temperatures are preferably of a high order of magnitude in the range of 1,000F. up to about l,300 or 1400F. However, it is contemplated under some circumstances employing a lower temperature for some specific applications. Generally the temperature will not be below 900F. Operating pressures may be in the range of from about atmospheric pressure up to several atmospheres depending upon the system employed. Generally the pressure will be selected from within the range of from about psig. up to about 50 or 75 psig. Operating space velocities are relatively high and are selected to provide catalyst-oil contact times in the virgin feed cracking step in the range of from about 1 second up to about 12 seconds and more usually in the range of 4 to 8 seconds. Contact times used in the recycle cracking step will be in the range of about 2 to 15 seconds and may be a few seconds longerthan employed for of virgin feed cracking step. On the other hand, the catalyst to oil ratio used for converting virgin feed will be from about 4 to 10 to l and from about 6 to to l for converting the recycle and coker feed material.
The method of the present invention is particularly suitable for effecting high temperature, above 1000F., catalytic cracking operations which have been found to yield products of superior quality and quantity relative to that attainable by more convetional operations now relied upon. Thus the concepts discussed herein are particularly directed and designed for high temperature, short contact time cracking operations in the alyst composition comprising two distinct cracking components. One component is a catalytically active form of a ZSM-S type zeolite. The other component may be any other suitable zeolite catalytic cracking component which is active with respect to converting the molecular constituents of the hydrocarbon charge and products of cracking to desired components. Generally speaking, the cracking component may be a porous cracking component such as silica/alumina and more particularly a crystalline aluminosilicate zeolite cracking component having relatively uniform pore dimensions and a pore size selected from within the range of 6 to 15 Angstrom units. These relatively large pore zeolite components will admit both normal and isoaliphatics. Particularly desirable zeolites include the synthetic faujasites known as zeolite X and zeolite Y. Other large pore zeolites may also be employed. The weight ratio of the porous siliceous cracking component, (e.g., synthetic faujasitelto the ZSM-S component may be selected from between 0.1 and 20.
The ZSM-5 type zeolites which may be used in combination with relatively large pore zeolites in the novel cracking process of this invention are considered to be of a relatively intermediate pore size. Thus, the ZSM-S type catalysts used in the novel process of this invention will allow entry into their internal pore structure or normal aliphatic compounds and slightly branched aliphatic compounds, particularly monomethyl substituted compounds, yet substantially exclude all compounds containing at least a quaternary carbon atom or having a molecular dimension equal to or substantially greater than a quaternary carbon atom. Additionally, aromatic compounds having side chains similar to the normal aliphatic compounds and slightly branched aliphatic compounds above described could have side chains enter the internal pore structure of the instant catalysts. Thus, if one were to measure the selectivity of the ZSM-S type materials employed in the process of this invention with regard to their ability to sorb nhexane in admixture with 2-methyl pentane, i.e., the ability to selectively sorb hexane from a mixture of the same with isohexane, these catalysts would have to be stated as being non-shape selective. It should be immediately apparent, however, that the term selectivity has a far greater significance than merely the ability to preferentially distinguish between nonnal paraffins and isoparaffins. Selectivity on shape is theoretically possible at any shape or size although, quite obviously, such selectivity might not result in an advantageous catalyst for any and all hydrocarbon conversion processes.
While not wishing to be bound by any theory of operation nevertheless, it appears that the crystalline zeolitic materials of the ZSM-5 type employed in the instant invention cannot simply be characterized by the recitation of a pore size or a range of pore sizes. lt appears that the uniform pore openings of this new type of zeolite are not approximately circular in nature, as is usually the case in the heretofore employed zeolites, but rather, are approximately elliptical in nature. Thus, the pore openings of the instant zeolitic materials have both a major and a minor axes, and it is for this reason that the unusual and novel molecular sieving effects are achieved. This elliptical shape can be referred to as a keyhole. It appears that the minor axis of the elliptical pores in the zeolites apparently have an effective size of about 5.5 Angstrom units. The major axis appears to be somewhere between 6 and about 9 Angstrom units. The unique keyhole molecular sieving action of these materials is presumably due to the presence of these approximately elliptically shaped windows controlling access to the internal crystalline pore structure.
A test method has been devised in order to determine whether or not a zeolite possess the unique molecular sieving properties desired to be combined with alarge pore zeolite in order to carry out the novel conversion process of this invention. In the test method a candidate zeolite free from any matrix or binder is initially converted to the so-called acid or hydrogen form. This procedure involves exhaustive exchange with an ammonium chloride solution in order to replace substantially all metallic cations originally present. The sample is then dried, sized to 2030 mesh and calcined in air for 16 hours at 550F. One gram of the so-treated zeolite is then contacted with benzene at a pressure of 12 mmHg. at a temperature of 25C. for a time period of 2 hours. Another gram sample is contacted with mesitylene at a pressure of 15 mmI-lg. at a temperature of 25C for a period of 6 hours. A preferred zeolite is one whose acid form will adsorb at least 3.0 weight percent benzene and less than 1.5 weight percent mesitylene at the above-recited conditions.
Examples of zeolitic materials which are operable in the process of this invention are ZSM-S type which family includes not only ZSM-S but also ZSM-8 zeolites. ZSM-S type materials are disclosed and claimed in copending application Ser. No. 865,472, filed October 10, 1969, now U.S. Pat. No. 3,702,886 and ZSM-8 is disclosed and claimed in copending application Ser. No. 865,418 filed Oct. 10, 1969 now abandoned. A process utilizing a combination of ZSM-5 type zeolites and large pore zeolites is disclosed in Ser. No. 78,573 filed Oct. 6, 1970.
The family of ZSM-5 compositions has the character- I istic X-ray diffraction pattern set forth in Table 1, hereinbelow. ZSM-S compositions can also be identified, in terms of mole ratios of oxides, as follows:
i Mgmo I M203 YO ZHzO wherein M is a cation, n is the valence of said cation, W is selected from the group consisting of aluminum and gallium, Y is selected from the group consisting of silicon and germanium, and z is from O to 40.'1n a preferred synthesized form, the zeolite has a formula, in terms of mole ratios of oxides, as follows:
0.9 i 0.2 M O A1 -100 SiO 2 H 0 and M is selected from the group consisting of a mixture of alkali metal cations, especially sodium, and tetraalkylammonium cations, the alkyl groups of which preferably contain 2-5 carbon atoms.
In a preferred embodiment of ZSM-5, W is aluminum, Y is silicon and the silica/alumina mole ratio is at least 10 and ranges up to about 60.
Members of the family of ZSM-5 zeolites possess a definite distinguishing crystalline structure whose X-ray diffraction pattern shows the following significant lines:
TABLE 1 interplanar Spacing d(A) Relative Intensity These values as well as all other X-ray data were determined by standard techniques. The radiation was the K-alpha doublet of copper, and scintillation counter spectrometer with a strip chart pen recorder was used.
The peak heights, 1, and the positions as a function of 2 times theta, where theta is the Bragg angle, were read from the spectrometer chart. From these the relative intensities, I/l where 1,, is the intensity of the strongest line or peak, and d (obs.), the interplanar spacing in A, corresponding to the recorded lines, were calculated. In Table l the relative intensities are given in terms of the symbols S strong, M medium, MS medium strong, MW medium weak and VS very strong. It should be understood that this X-ray diffraction pattern is characteristic of all the species of ZSM-5 compositions. Ion exchange of the sodium ion with cations reveals substantially the same pattern with some minor shifts in interplanar spacing and variation in relative intensity. Other minor variations can occur depending on the silicon to aluminum ratio of the particular sample, as well as if it has been subjected to thermal treatment. Various cation. exchanged forms of ZSM-5 have been prepared. X-ray powder diffraction patterns of several of these forms are set forth below. The ZSM-S forms set forth below are all aluminosilicates.
TABLE 2 X-Ray Diffraction ZSM-S Powder in Cation Exchanged Forms d Spacing Observed CaCl TABLE 2-Continued Made HCl NaCl CaCl RECl AgNO Zeolite ZSM-S can be suitably prepared by preparing solution containing tetrapropyl ammonium hydroxide, sodium oxide, an oxideof aluminum or gallium, an oxideof silica and water and having a composition, in terms of mole ratios of oxides, falling within the following ranges:
TABLE 3 Particularly Broad Preferred Preferred OH'lSiO 0.07-1.0 0.1-0.8 0.2-0.75 R N-l-/(R. N +Na*) 0.2-0.95 0.3-0.9 0.4-0.9 H O/OH 10-300 10-300 10-300 wherein R is propyl, W is aluminum and Y is silicon maintaining the mixture until crystals of the zeolite are formed. Thereafter the crystals are separated from the liquid and recovered. Typical reaction conditions consist of heating the foregoing reaction mixture to a temperature of from about C to 175C for a period of time of from about six hours to 60 days. A more preferred temperature range is from about to C with the amount of time at a temperature in such range being from about 12 hours to 20 days.
. The digestion of the gel particles is carried out until crystals form. The solid product is separated from the reaction medium, as by cooling the whole to room temperature, filtering, and water washing.
ZSM-S is preferably formed as an aluminosilicate. The composition can be prepared utilizing materials which supply the appropriate oxide. Such compositions include for "an aluminosilicate, sodium aluminate, alumina, sodium silicate, silica hydrosol, silica gel, silicic acid, sodium hydroxide and tetrapropylammonium hydroxide. It will be understood that each oxide component utilized in the reaction mixture for preparing a member of the ZSM-S family can be supplied by one or more initial reactants and they can be mixed together in any order. For example, sodium oxide can be supplied by an aqueous solution of sodium hydroxide, or by an aqueous solution of sodium silicate; tetrapropylammonium cation can be supplied by the bromide salt. The reaction mixture can be prepared either batchwise or continuously. Crystal size and crystallization time of the ZSM-S composition will vary with the nature of the reaction mixture employed. ZSM-8 can also be identified, in terms of mole ratios of oxides, as follows:
M 1 0 I A1203 Z H2O wherein M is at least one cation, n is the valence thereof and z is from 0 to 40. In a preferred synthesized form, the zeolite has a formula, in terms of mole ratios of oxides, as follows:
i Mg o I A1203 2 H20 TABLE 4 dA m. 1/1, dA
TABLE 4-Cont1nued dA 1/1, 1/1, dA
9.0 6 l 2.78 7.42 10 4 2.73 7.06 7 1 2.68 6.69 5 3 2.61 6.35 12 l 2.57 6.04 6 l 2.55 5.97 12 l 2.51 5.69 9 6 2.49 5.56 13 1 2.45 5.36 3 2 2.47 5,12 4 3 2.39 5.01 7 1 2.35 4.60 7 1 2.32 4.45 3 1 2.28 4.35 7 1 2.23 4.25 18 1 2.20 4.07 20 l 2.l7 4.00 1 2.12 3.85 100 1 2.11 3.82 57 1 2.08 3.75 25 1 2.06 3.71 30 6 2.01 3.64 26 6 1.99 3.59 2 2 1.95 3.47 6 2 1.91 3.43 9 3 1.87 3.39 5 1 1.84 3.34 18 2 1.82 3.31 8
Zeolite ZSM-8 can be suitably prepared by reacting a solution containing either tetraethylammonium hydroxide or tetraethylammonium bromide together with sodium oxide, aluminum oxide, and an oxide of silica and water.
The relative operable proportions of the various ingredients have not been fully determined and it is to be immediately understood that not any and all proportions of reactants will operate to produce the desired zeolite. In fact, completely different zeolites can be prepared utilizing the same starting materials depending upon their relative concentration and reaction conditions as is set forth in US. Pat. No. 3,308,069. In general, however, it has been found that when tetraethylammonium hydroxide is employed, ZSM-8 can be prepared from said hydroxide, sodium oxide, aluminum oxide, silica and water by reacting said materials in such proportions that the forming solution has a composition in terms of mole ratios of oxides falling within the following range SiO /Al O from about 10 to about 200 Na O/tetraethylammonium hydroxide from about 0.05 to 0.20
Tetraethylammonium hydroxide/SiO from about 0.08 to 1.0
H oltetraethylammonium hydroxide from about 80 to about 200 Thereafter, the crystals are separated from the liquid and recovered. Typical reaction conditions consist of heating the foregoing reaction mixture to a temperature of from about 100C to 175C for a period of time of from about 6 hours to 60 days. A more preferred temperature range is from about 150 to 175C with the amount of time at a temperature in such range being from about 12 hours to 8 days.
The foregoing product is dried, e.g., at 230F. for from about 8 to 24 hours. Of course, milder conditions may be employed if desired, e.g., room temperature under vacuum.
As has heretofore been stated, a zeolite of the ZSM-5 type above-described is used in conjunction with a large pore zeolite, i.e., one having a pore size greater than 7 Angstrom units which has the ability to act upon substantially all the components usually found in a commercial gas oil. Large pore aluminosilicates of this type are well known and include natural and synthetic faujasite of both the X and Y type, as well as zeolite L.
Of these materials, zeolite Y is particularly preferred.
Both the large pore zeolites and-the ZSM-5 type zeolites used in the instant invention usually have the original cations associated therewith replaced by a wide variety of other cations according to techniques well known in the art. Typical replacing cations would include hydrogen, ammonium and metal cations including mixtures of the same. Of the replacing metallic'cations, particular preference is given to cations of rare earth, Mg, Zn, Mn, Alt, and Ca.
Typical ion exchange techniques would be to contact the particular zeolite with a salt of the desired replacing cation or cations. Although a wide variety of salts can be employed, particular preference is given of chlo-. rides, nitrates and sulfates.
Representative ion exchange techniques are disclosed in a wide variety of patents including US. Pat. Nos. 3,140,249; 3,140,251; and 3,140,253.
Following contact with the salt solution of the desired replacing cation, the zeolites may be washed with water and dried at a temperature ranging from F. to about 600F. and thereafter heated in air or other inert gas at temperatures ranging from about 500 to 1500F for periods of time ranging from 1 to 48 hours or more. It has been further found in accordance with the invention that catalysts of improved selectivity and having other beneficial properties in catalytic cracking are ob-' tained by subjecting the zeolite to treatment with steam at elevated temperatures ranging from 800 to 1,600F
and preferably 1,000F and 1,500F. The treatment may be accomplished in atmospheres consisting partially or entirely of steam. This treatment may be accomplished within a commercial cracking unit, e.g., by gradual addition of the unsteamed catalyst to the unit.
A similar treatment can be accomplished at lower temperatures and elevated pressures, e.g. 350700F at 10 to about 200 atmospheres.
The novel catalyst composites of this invention, in a particular embodiment, comprise a physical mixture of at least two different cracking components, one being an aluminosilicate having a pore size greater than about 7 Angstrom units. In one embodiment, a mixture of catalyst particles is used in'which each particle contains only one of the two types of zeolites. Thus, for example, a mixture of spray dried particles comprising ZSM-S type crystals in a matrix and particles comprising faujasite crystals in a matrix may be added as make-up to the cracking unit. Alternatively, the catalyst components may be pelleted, cast, molded, spray-dried or otherwise formed into pieces of desired size and shape such as rods, spheres, pellets, etc.
The compositing of the aluminosilicate with an inorganic oxide can be achieved by several methods wherein the aluminosilicates are reduced to a particle size less than 40 microns, preferably less than 10 microns, and intimately admixed with an inorganic oxide while the latter is in a hydrous state such as in the'form of hydrosol, hydrogel, wet gelatinous precipitate, or in a dried state, or a mixture thereof. Thus, finely divided aluminosilicates can be mixed directly with a siliceous gel formed by hydrolyzing a basic solution of alkali metal silicate with an acid such as hydrochloric, sulfuric, acetic, etc. The mixing of the three components can be accomplished in any desired manner, such as in a ball mill or other types of mills. The aluminosilicates also may be dispersed in a hydrosol obtained by reacting an alkali metal silicate with an acid or alkaline coagulant. The hydrosol is then permitted to set in mass to a hydrogel which is thereafter dried and broken into pieces of desired shape or dried by conventional spray drying techniques or dispersed through a nozzle into a bath of oil or other water-immiscible suspending medium to obtain spheroidally shaped bea particles of catalyst such as described in US. Pat. No. 2,384,946. The aluminosilicate siliceous gel thus obtained is washed free of soluble salts and thereafter dried and/or calcined as desired.
In a like manner, the aluminosilicates may be incorporated with an aluminiferous oxide. Such gels and hydrous oxides are well known in the art and may be prepared, for example, by adding ammonium hydroxide, ammonium carbonate, etc. to a salt of aluminum, such aluminum chloride, aluminum sulfate, aluminum .nitrate, etc., in an amount sufficient to form aluminum hydroxide which, upon drying, is converted to alumina. The aluminosilicate may be incorporated with the aluminiferous oxide while the latter is in the form of hydrosol, hydrogel, or wet gelatinous precipitate or hydrous oxide, or in the dried state.
The catalytieally inorganicv oxide matrix may also consist of a plural gel comprising 1 a predominant amount of silica with one or more metals or oxides thereof selected from Groups 1B, 11, III, IV, V, VI, VII, and VIII of the Periodic Table. Particular preference is given to plural gels of silica with metal oxides of Groups llA, Ill and Na of the Periodic Table, especially wherein the metal oxide is rare earth oxide, magnesia, alumina, zirconia, titania, beryllia, thoria, or combination thereof. The preparation of plural gels is well known and generally involves either separate precipitation or coprecipitation techniques, in which a suitable salt of the metal oxide is added to an alkali metal silicate and an acid or base, as required, is added to precipitate the corresponding oxide. The silica content of the siliceous gel matrix contemplated herein is generally within the range of 55 to 100 weight percent with the metal oxide content ranging from 0 to 45 percent.
from about 500F. to 1,600F. for periods of time rang-,
ing from 1 to 48 hours or more. It is to be understood that the aluminosilicate can also be calcined prior to incorporation into the inorganic oxide gel. It is also to be understood that the aluminosilicate or aluminosilicates need not be ion exchanged prior to incorporation in a matrix but can be so treated during and/or after incorporation into the matrix. Preferably, the zeolite is metal exchanged, calcined and thereafter given a second exchange with a metal or hydrogen precursor.
It has been further found in accordance with the invention that catalysts of improved selectivity and having other beneficial properties in gas oil cracking are obtained by subjecting the catalyst product to a mild steam treatment carried out at elevated temperatures of 800F. to 1,600F. and preferably at temperatures of about 1,000F. to 1,500F. The treatment may be accomplished in an atmosphere of 100 percent steam or in an atmosphere consisting of steam and air or a gas which is not harmful to the aluminosilicate. The steam treatment apparently provides beneficial properties in the aluminosilicate compositions and can be conducted before, after or in place of the calcination treatment.
The particle size of each type of zeolite making up the catalyst system is not narrowly critical but should be less than 100 microns and particle sizes within the range of from less than 0.1 to 10 microns are preferred. It is also to be noted that each individual component in the catalyst system need not be of the same particle size.
The particular proportion of one component to the other in the catalyst system is also not narrowly critical and can vary over an extremely wide range. However, it has been found that for most purposes the weight ratio of the ZSM-S type aluminosilicate to the larger pore size aluminosilicate with which it is mixed can range from 0.05:1 up to 10:1 and preferably from 1:3 up to 2:1 and still more preferably 1:2 to 1:1.
The ZSM-S type crystalline aluminosilicates and the crystalline aluminosilicates with pores greater than 7 Angstroms may be added to a cracking unit as a mixture of crystallites within the same particles of catalyst composite, whether the particles are beads, extrudates, or spray-dried microspheres. Alternatively, a mixture of particles of fluidizable particle size may be added to the cracking unit, some particles containing only the ZSM-S type aluminosilicate crystallites and the other particles containing only the large pore aluminosilicate crystallites. In either case, the ratio of ZSM-5 type aluminosilicates to large pore aluminosilicates should be within the range of 1:20 to 10:1. The ratio of aluminosilicates within this range is controlled to produce the most desirable balance of high octane gasoline and C and C olefin yields.
Within the above description of the aluminosilicates which can be physically admixed in a porous matrix to prepare the catalysts of this invention, it has been found that certain aluminosilicates provide superior results when employed in catalytic cracking operations.
First of all, it is preferred that there be a limited amount of alkali metal cations associated with the aluminosilicates since the presence of alkali metals tends to suppress or limit catalytic properties, the activity of cations of monovalent metals. Of the trivalent metal cations, the most preferred are rare earth metal'cations, either individually was a mixture of rare earth metal cations.
Additionally, it is particularly preferred to have at least some protons or proton precursors associated with the aluminosilicate.
It is also preferred that both the aluminosilicates have an atomic ratio of silicon to aluminum of at least 1.25 preferably 1.8 and even more desirably at least 2.0.
It is to be understood, however, that this invention includes the use of catalyst compositions wherein both aluminosilicates are of the same class, e.g., both metal aluminosilicates; of different classes, e.g., one metal and one acid aluminosilicate; in the same matrix or in different matrixes, i.e., one aluminosilicate in silicaalumina and the other in silica-zirconia.
In the process of the present invention, the cracking catalyst has a particle size such that it can be passed in fluid flow through the risers, catalyst separators, stripper, transfer conduits and the regenerator. The particle size will generally be between and 100 microns in diameter, preferably 40 to 80 microns. A particle size of about 60 microns diameter is considered optimum.
Low coke producing crystalline aluminosilicate cracking catalysts of the type which may be used with particular advantage in the method and processing schemes of this invention are more particularly exemplified by the following table and FIGS. 1 and 2 presented herewith. The table provides the chemical properties of several crystalline zeolite containing cracking catalysts having different selectivity characteristics and particularly coke producing characteristics.
coke' on charge versus conversion after steam treatmentof identified catalyst compositions.
7 FIG. 3 diagrammatically represents one arrangementof processing steps comprising a dual riser operation for catalytic conversion of hydrocarbon reactant and interconnecting catalyst transfer means for conveying catalyst particles to the regenerator and from the regenerator to the riser conversion operations.
FIG. 4 diagrammatically shows a modified embodiment of FIG. 3 wherein the dual riser catalyst collection hopper is separated by a vertical baffle provided with catalyst flow through passageways for passing catalyst in either direction through the vertical baffle.
FIG. 5 diagrammatically shows -a further modification of FIG. 3 comprising a common fluid bed of catalyst discharged from each riser and catalyst withdrawal means above the stripping section for providing the catalyst to the second riser reactor.
DISCUSSION OF SPECIFIC EMBODIMENTS The operating conditions employed in obtaining the data plotted in-FIGS. l and 2 are identified on the Fig ther steam or thermal treatment. Close examination of the data shows that several of the catalyst compositions are much higher coke producers after thermal treat-' ment. The thermally treated catalyst is designed to rep- TABLE 5 CHEMICAL PROPERTIES OF CRYSTALLINE ZEOLITE CRACKING CATALYSTS Sieve Chemical Analysis, Wt Loss ned 3 hrl1200FlAir Mean Cata- Type SiO A1 0 RE O Na O ZrO 80., on Surface Pore Packed Dialyst and lgni- Area, Vol Density, meter Level tion mlg cc/g g/cc Micron A 7.5% REX 84.4 13.8 1.84 0.05 415 0.93 0.47 77 B 5% REY 85.0 14.0 0.96 0.04 13-15 442 0.89 0.46
C 10% REY 77.4 18.0 2.5 0112 2.0 0.2 12-14 328 0.67 0.56
D 1020% HY 85.5 14.0 0.05 486 0.79
E 10% REX 66.0 31.0 2.9 0.09 280 0.58 0.63 79 F 12% REX 64.6 32.6 2.8 0.04 269 0.69 0.55 70 G 1214% REX 58.7 35.0 4.5 0.07 0.12 15.4 270 0.66
1-1 15% REY 66.6 29.3 3.6 0.46 329 0.58 0.58
I 15% REY 60.5 36.8 2.5 0.22 203 0.52 0.65 71 .1 3% REX 73.0 25.7 0.87 0.04 295 0.55 0.61
K 6.7% REX 64.0 33.0 1.67 0.11 227 0.57 0.64
L 5% REY 65.7 33.2 1.00 0.05 260 0.55 0.64
M 45% REY 67.5 31.7 0.66 0.14
N 7-8% REY 82.4 15.2 2.30 0.07 0.68 0.46
0 11-12% REX? 51.7 45.5 2.20 0.62 169 0.38 0.79 51 P 10% REY 53.4 43.8 2.01 0.77
Q 5% REY 84.8 14.0 1.18 0.05 0.18 14 561 0.72 0.54
R 10% REY 71.1 26.7 2.06 0.09
By difference on a basis. Crystallinity shows 30% shift toward X-type material. (3) Crystallinity indicates sieve of an X-Y type.
BRIEF DESCRIPTION OF THE DRAWINGS resent a freshly prepared catalyst. The steam treated catalyst composition is designed to represent equilib- FIG. 1 is a plot of data obtained, Coke on charge ver- 65 rium catalyst existing in a commercial operation. It is sus conversion after thermal treatment of the identified catalyst compositions.
FIG. 2, on the other hand, is a plot of data obtained,
to be particularly observed that catalyst compositions I, H, B and F form substantially a straight line on FIG.
1 and are representative of a higher coke making catalyst than that experienced with catalyst C represented by the curve on the figure.
The above observation is also made with respect to the data plotted on FIG. 2. Thus, it is directionally clear from these informations just how the coke on charge varies with conversion and for different catalyst compositions. Furthermore, it is clear that merely reciting that a cracking catalyst comprises a crystalline zeolite as a cracking component is insufficient to identify its activity and/or selectivity or the operating parameters in which it is most desirably employed.
It is also evident from these data of the need for highly sophisticated catalyst conversion systems such as that described herewith which will permit one to take advantage of a particular catalyst composition in an optimum manner. This does not mean to say, however, that other catalyst compositions such as those producing more coke than the low coke producing catalyst C of this invention cannot be used with a considerably greater efficiency in the method and system of this invention. On the contrary such catalyst may also be employed with a high degree of sufficiency and in many instances at a greater efficiency than permitted in now available in fluid catalyst systems.
Referring now to FIGS. 3, 4 and by way of example, there is shown schematically arrangements of catalyst systems for practicing the processing concepts herein described and particularly going to the essence of this invention.
Referring now specifically to FIG. 3, there is sho diagrammatically an arrangement of means for providing plural stages of riser cracking, separation of the riser effluent into a hydrocarbon phase and a catalyst phase, stripping of the separated catalyst, passing one of the stripped catalyst phases to regeneration, and
using part or all of the other of these catalyst phases in combination with freshly regenerated catalyst in one of the risers provided for converting heavy hydrocarbons such as high molecular weight hydrocarbons, residual hydrocarbons, coker gas oils, or recycle feed materials either alone or combined with other materials difficult to crack and producing high levels of coke. In the arrangement of FIG. 3, a hydrocarbon feed such as a virgin feed, fresh feed or fractions thereof is introduced to the processing arrangement by conduit 2 to furnace 4 wherein the feed is preheated to a desired elevated temperature within the range of 300F. to 800F. or higher. Preheating of the hydrocarbon feed may occur in the presence of or absence of a suitable diluent gasiform material described herein in such quantities as to effect a desired reduction in the hydrocarbon partial pressure. The diluent employed may be steam, dry hydrocarbon gases or gasiform material comprising C hydrocarbons and particularly C hydrocarbons. The amount of diluent employed will, of course, vary with the type and boiling range of the fresh feed charge employed and will be used in a lesser amount when converting, for example, a low molecular weight hydrocarbon feed as distinguished from a virgin feed material boiling up to about l,000F. or l,200F. The preheated feed with combined diluent at a desired elevated temperature in conduit 6 is combined with freshly regenerated catalyst withdrawn from the regenerator by provided standpipe 8 to form a suspension having a catalyst to oil ratio selected from within the range of 4 to 10/ l at a temperature of at least about 900F. and preferably at a temperature in the range of l,00OF. up to about l,200F. The temperature selected for accomplishing conversion in the riser will, of course, depend upon the specific product desired from a given hydrocarbon charge material. For example, a product com prising significant amounts of olefins and/or isobutane may be had with very little coke production particularly when employing high temperatures and the low coke producing crystalline zeolite cracking catalyst such as catalyst C discussed hereinbefore used alone or in combination with the ZSM-S type of catalyst. Operating conditions will be employed which generally maximize the yield of gasoline boiling range product. Under some conditions it may be desirable to produce significant amounts of olefins suitable for use in alkylation reactions to produce gasoline boiling range product in which case the low coke producing catalyst C would be used alone in the absence of ZSM-5 type of catalyst particles. In any event, the suspension formed, as described above, is caused to move through riser 10 at a velocity designed to provide a residence time therein selected from within the range from 1 to 12 seconds and preferably restricted to within the range of 4 to 8 seconds. Thus, when a very short residence time in the riser is desired, for example, in the rangeof l to 4 seconds, it is contemplated relying upon the gasifonn diluent material to form a suspension with freshly regenerated catalyst which suspension is caused to flow through an initial portion of the riser reactor before bringing the hydrocarbon reactant material in contact therewith in a downstream portion of the reactor. Thus the fresh feed riser may have a plurality of spaced apart hydrocarbon feed inlets 7 and .7 throughout the length thereof so as to facilitate varying the residence time that the hydrocarbon reactant is in contact with the catalyst suspension in the riser reactor. In any event the hydrocarbon charge isconverted in the riser for a predetermined residence time before being separated from suspended catalyst by discharge into a hopper 14 wherein the separation is facilitated by one or more cyclone separator positioned in hopper 14. It is preferred that separation of the suspension be initially completed in one or more cyclone separators positioned within or external to the catalyst collecting hopper to avoid undesired prolonged cracking. Thus discharging the suspension into the dilute catalyst phase of a hopper is not sufficient to obtain the separation desired by this invention. It is also important to the method and concept of this invention that the catalyst separated from the suspension discharged from the riser be recovered at its highest temperature so that the available heat of the catalyst can be utilized to maximum advantage as hereinafter provided. Thus quenching of the suspension or separated catalyst should be avoided.
In the arrangement of FIG. 3, the suspension discharged from the fresh feed riser is separated as by cyclone separator 11 with the separated hydrocarbons being passed to a collection chamber 12 in the upper portion of the catalyst hopper 14. The collected hydrocarbon vapors are passed from chamber 12 by conduit 16 to suitable fractionation equipment not shown but conventional in the art.
The catalyst separated from the suspension is passed from the cyclone by diplet 18 to a dense fluid bed of catalyst particles 20 therebelow. Stripping gas such as steam is introduced at an elevated stripping temperature to the lower portion of the catalyst bed 20 by conduit 22 so as to flow upwardly therethrough without significant cooling thereof to remove entrained vaporous hydrocarbons from the hot catalyst particles. Provisions are made for passing stripped hydrocarbons and stripping gas through, for example, a separate cyclone separator before the stripping gasiform material passes into the collection chamber 12. Any entrained catalyst fines separated in such a cyclone separator are returned to the catalyst bed by a suitably provided dipleg. A substantially vertical baffle 24 is provided in the hopper and extending upwardly from the bottom thereof for keeping the catalyst separated from riser l collected as a separate bed of catalyst in the lower portion of the hopper. The upper portion of the hopper and above the upper dense phase of the separated catalyst beds may be in open communication with one another. The catalyst recovered as catalyst bed and being at a relatively high temperature not substantially below the riser discharge temperature also contains a relatively small amount of deposited carbonaceous material thereon and therefore has a considerable amount of residual activity and heat available for further use, as herein described. The catalyst in bed 20 is thus withdrawn from the lower portion thereof by standpipe 26. The catalyst withdrawn by standpipe 26 is mixed with hot freshly regenerated catalyst withdrawn from regenerator 50 by standpipe 28. The amount of freshly regenerated catalyst combined with partially spent catalyst will vary depending upon the temperature of the partially spent catalyst and the temperature selected for converting hydrocarbon feed in the second riser. Thus the temperature of the catalyst mixture will depend upon the conversion temperature desired in the second riser reactor generally above about 900F. for processing heavy charge materials such as recycle hydrocarbons, coker gas oils and/or high molecular weight charge materials either alone or in admixture with one another. In any event the hot catalyst mixture comprising partially used catalyst and freshly regenerated catalyst is at a temperature sufficient to supply a major portion of the endothermic heat requirements of the recycle conversion riser. In some instances the amount of regenerated catalyst combined with partially spent catalyst will be sufficient to make up for a substantial portion of not all of the endothermic heat loss of the partially spent catalyst and this heat makeup in combination with the hydrocarbon charge preheat will be sufficient to effect the elevated temperature conversion desired in the second riser reactor herein referred to as the recycle riser reactor. Thus the second riser reactor confines a moving suspension of catalysthydrocarbon charge and diluent gasiform material having a temperature selected from within the range of 950F. up to about 1050F. and a catalyst to oil ratio in the range of 6 to /1. The suspension is caused to move through the second riser reactor for a predetermined residence time selected from within the range of 2 to 15 seconds. It is contemplated, as hereinbefore dis- .37" and 37' so as to provide contact times of about 2'seconds and higher. Thus the plurality of hydrocarbon feed inlets spaced along the length to the recycle riser facilitate obtaining very low hydrocarbon resifeed is charged by conduit 32 to fumace34 wherein it is heated to an elevated cracking temperature either in the presence of or absence of diluent gasiform material. The amount of gasiform diluent material employed will depend, as hereinbefore discussed, upon the hydrocarbon charge employed and the extent to which it is desired to lower the partial pressure of the hydrocarbon charge to be converted. In any event cracking of a recycle hydrocarbon charge is enhanced by relatively low hydrocarbon partial pressures and sufficient diluents should be employed to optimize this effect. The charge preheated to an elevated temperature and combined with the catalyst mixture as above discussed is then passed through recycle riser reactor 36 at a temperature selected from within the range of 900F. up to about 1,200F. The suspension undergoes conversion of the hydrocarbon charge during passage through the riser reactor before being separated as by discharge into hopper 14 containing cyclone separator equip-- ment or by direct discharge into a cyclone separator 38 attached to the end of the riser and located within hopper 14. The suspension in riser 36 is separated in a specific example in cyclone separator 38 into a hydrocarbon vapor phase and a catalyst particle phase. The hydrocarbon vapors are collected in chamber 12 and thence passed to a fractionator not shown. The separated catalyst is passed by a suitable dipleg to a dense fluid catalyst bed 40. The dense fluid bed of catalyst particles 40 are strippedof entrained hydrocarbon vapors by steam or other suitable stripping medium introduced to the lower portion of the bed by conduit 42. The stripped catalyst comprising deposited carbonaceous material is removed by standpipe 44 for passage to catalyst regeneration.
In the arrangement of FIG. 3, the catalyst in standpipe 44 is combined with regeneration gas such as air or other suitable oxygen containing regeneration gas mixture introduced by conduit 46 to form a suspension which is then passed through the riser regenerator 48 for discharge in the upper dispersed phase of regeneration vessel 50. This riserv regenerator may also discharge in the upper or intermediate portion of the dense bed of catalyst in the regenerator. The catalyst in riser regenerator 48 undergoes partial regeneration therein and upon discharge, for example, into the dilute phase is separated as by settling and becomes a part of the dense bed of catalyst therebelow. It is important to recover all available heat generated in the catalyst phases and flue gases. Regeneration gas such as air or oxygen supplemented gasiform material is introduced to the bottom or lower portion of catalyst bed 52 by inlet means 56. Gaseous products of combustion or flue gases pass through one or more catalyst cycle separator not shown in the upper portion of the regenerator where entrained catalyst fines are separated from the flue gas before the flue gases are removed from the upper portion of the regenerator by conduit 54.
FIGS. 4 and 5 differ from the processing arrangement of FIG. 3 disussed above primarily in the catalyst hopper arrangement or design housing the cyclone separators at the discharge of each riser and relied upon for cascading catalyst particles separated from each riser effluent to the separate catalyst streams removed therefrom.
Specifically FIG. 4 departs from FIG. 3 by providing a vertical baffle 24 similar to that described in FIG. 3 and containing catalyst flow through slots 60. In this arrangement transfer of catalysts between that discharged from riser and forming catalyst bed and that discharged from riser 36 forming catalyst bed 40 is provided so as to provide a further control on the catalyst/oil ratio particularly desired in riser 36. For example, in an operation where higher temperatures of the catalyst passed to the regenerator are desired the catalyst discharged from fresh feed riser 10 may be caused to flow from bed 20 through the slots in the vertical partition and be combined with the more contaminated catalyst in bed 40 discharged from riser 36 and recycled thereafter to the regenerator. Of course the reverse flow of catalyst through slots 60 may be had by changing the upper bed level of catalyst bed 40 discharged from riser 36 to be above the upper bed level of catalyst 20 on the other side of baffle 24 and dis charged from riser 10. This latter catalyst flow would be used in the case where a high catalyst to oil ratio is desired in riser 36 in addition to increasing the level of carbonaceous material on the catalyst eventually passed to the regeneration zone.
In the arrangement of FIG. 4, catalyst withdrawn by standpipe 26 is normally combined with freshly regenerated catalyst and the thus formed catalyst mixture is employed in riser 36 similarly to that discussed with respect to FIGS. 3 and 4. In addition, catalyst withdrawn by standpipe 44 is passed to the regenerator in the manner similar to that discussed with respect to FIGS. 3 and 4. The variation in operation of FIG. 4 over that of FIG. 3 resides primarily in providing catalyst flow between the separated catalyst beds 20 and 40 respectively and as desired.
FIG. 5, on the other hand, is a further variation on the concept of FIG. 4 in that the catalyst discharged from each riser is accumulated as a common dense fluid bed of catalyst which is stripped in the lower portion thereof by stripping gas in a relatively confined stripping section provided with alternately staggered baffling to cause intimate contact between catalyst particles and stripping medium.
In the arrangement of FIG. 5, catalyst is withdrawn by standpipe 44 for transfer to the regenerator as discussed above. In this arrangement catalyst withdrawal standpipe 26 is extended upwardly into the dense fluid bed of catalyst particles to a catalyst withdrawal cone 62 positioned above the relatively confined catalyst stripping section. Stripping steam is introduced to the lower portion of the bed of catalyst and will effect considerably stripping of the catalyst in the upper portion of the bed before it enters withdrawal cone 62. It is contemplated adding in this arrangement hydrocarbon material such as recycle oil, slurry oil or other high mol weight hydrocarbon material to be partially converted at a point substantially intermediate the withdrawal cone and an upper portion of the catalyst stripping section. In this embodiment additional carbonaceous material for heat balance purposes will be laid down on the catalyst and prior to its return to the regenerator.
The arrangement of FIG. 5 may be further modified by withdrawing all catalysts from beneath the stripping gas inlet whether it is returned directly to the regenerator as by standpipe 44 or passed to riser 36 as a mixture of catalyst particles with freshly regenerated catalyst. The amount of carbonaceous material deposited on the catalyst may be varied considerably.
In yet a further'embodiment, hopper 14 of FIG. 5 may be used in a manner resembling an elutriator wherein catalyst particles of reduced density by virtue of carbonaceous deposits therein will be caused to lo cate inan upper portion of the dense fluid catalyst bed with the more dense catalyst particles because of a small amount of carbonaceous deposits being located in a lower portion of the dense fluid catalyst bed. Thus the more dense particles will be withdrawn from thelower portion of the catalyst bed for passage to riser 36 in admixture with freshly regenerated catalyst and the catalyst particles containing a higher level of carbonaceous material being withdrawn from an upper portion of the catalyst bed for return to the regenerator as by withdrawal standpipe 4-4.
It is clearly evident from the discussions hereinbefore presented that the processing schemes discussed and variations thereto are directed to and designed particularly for the purpose of maximizing utilization of the activity and selectivity characteristics of low coke forming crystalline zeolite catalyst compositions in the conversion of hydrocarbon charge materials. Furthermore the combinations improved upon the heat balance of a coke deficient operation in a manner which will minimize extraneous heat requirements of the process, for example, minimize the preheat requirements for the hydrocarbon feed and regeneration gas. Thus in any of the embodiments discussed hereinbefore it is contemplated contacting the catalyst to be passed to regeneration either before or after final stripping thereof with a heavy hydrocarbon fraction such as a residual oil fraction which will operate to lay down a further increment of carbonaceous material on the catalyst for the purpose of improving the overall heat balance of the processing arrangement. For example, slurry oil containing catalyst fines and recovered from the bottom of the fractionator may be combined with the catalyst to be passed to regeneration as a means for recovering catalyst fines and providing heat producing carbonaceous material on the catalyst particles.
Regeneration of the catalyst particles may be accomplished as briefly discussed hereinbefore in any one of several different arrangements of carbonaceous material burning sequences with oxygen containing gas which will be effective in removing contaminating car bonaceous deposits from the catalyst particles to restore substantially, if not completely, the activity and selectivity of the catalyst particles. Thus the catalyst may be regenerated in a plurality of sequentially arranged dilute-phase catalyst bed regeneration zones to which oxygen may be added in separate increments to avoid overheating of the catalyst during regeneration of the catalyst at a temperature selected within the range of from about I000F. to as highas l400F. On the other hand, the dilute phase regeneration zones may be employed in combination-with a more dense phaseregeneration step to accomplish the same purpose. In any event the regeneration sequence selected should be one which will recover the major portion of heat available to burning of carbonaceous material and its ultimate use in the'conversion processing steps of the arrangements discussed above.
Having thus provided a general discussion of the improved method and concepts of this invention and described specific examples in support thereof going to the essence of the invention, it is to be understood that no undue restrictions are to be imposed by reason thereof except as defined in the following claims.
1. In a dual riser hydrocarbon conversion operation employing fluidizable catalyst particles, the improved method of operation which comprises,
regenerating a catalyst mixture comprising a large pore crystalline zeolite of the faujasite type in admixture with from to about 90 weight percent of a smaller pore crystalline zeolite of the ZSM-S type under conditions to heat the catalyst mixture to an elevated temperature in the range of 1,000F. up to about 1400F., passing heated catalyst to the inlet of a first riser conversion zone for admixture with gaseous materials such as dry gas and/or C and lighter gaseous hydrocarbons to form a first suspension therewith, passing said first suspension in admixture with a fresh gas oil feed upwardly through said first conversion zone under elevated temperature cracking conversion conditions of at least 1000F. and a hydrocarbon residence time less than about 12 seconds to cyclonic separation of the suspension into a hydrocarbon phase and a low coke containing catalyst phase, recovering the hydrocarbon phase and stripping the catalyst phase,
passing stripped catalyst mixture of relatively low coke level separated from said first conversion zone to the inlet of a second riser conversion zone in combination with an increment of the freshly regenerated catalyst,
forming a second suspension of catalyst with gaseous materials such as dry gas and/or C and lighter gas? eous hydrocarbons for passage upwardly through said second riser conversion zone,
passing said second suspension with higher coke producing hydrocarbon feed upwardly through said second conversion zone at a temperature of at least 950F. for a hydrocarbon residence time in the range of 2 to 15 seconds and a higher catalyst to oil ratio than employed in said first conversion zone,
cyclonically separating the second suspension at the discharge of the second conversion zone into a second hydrocarbon phase and a second catalyst phase,
recovering the second hydrocarbon phase and stripping the second catalyst phase of entrained hydrocarbons, and
passing stripped second phase catalyst to said catalyst regeneration.
2. The process of claim 1 wherein an amount of metal oxide which will enhance regeneration of the catalyst by converting carbon monoxide to carbon dioxide is admixed with the zeolite catalyst mixture and particularly the ZSM-S crystalline component when employed as separate .and discrete particles in admixture with particles of the large pore crystalline zeolite.
3. The method of claim 1 wherein catalyst discharged from each conversion zone is dumped into a common fluid bed of catalyst particles undergoing stripping with a stripping gas, a portion of the stripped catalyst is passed to catalyst regeneration and another portion of the stripped catalyst is combined with freshly regenerated catalyst for passage to the inlet of the second conversion zone.
4. The method of claim 1 wherein contaminated catalyst particles recovered from the second conversion zone stripping operation is regenerated in stages with at least one stage being a dense fluid bed regeneration operation.
5. The method of claim 1 wherein regeneration of the catalyst is accomplished in at least one dispersed phase regeneration operation discharging adjacent the upper interface of a dense fluid bed of catalyst being regenerated.
6. In a process for cracking hydrocarbons with crystalline zeolite cracking catalyst and regenerating coke deactivated catalyst wherein a plurality of riser conversion zones are relied upon for processing a first high boiling hydrocarbon feed in one conversion zone and a second high boiling hydrocarbon feed in another conversion zone and wherein catalyst separated from the product of the first conversion zone is passed in admixture with hot regenerated catalyst to the second conversion zone, the improvement which comprises,
using as the hydrocarbon conversion catalyst in the combination operation a mixture of a large pore low coke producing crystalline zeolite in combination with from 10 to weight percent of a crystalline zeolite of the ZSM-S type, forming a suspension of the catalyst mixture with gaseous hydrocarbons selected from the group comprising C and lighter dry gases and/or wet gases comprising C and lighter hydrocarbons in the inlet portion of each riser conversion zone, converting a first fresh gas oil feed in admixture with said suspension in the first conversion zone under operating conditions of temperature, catalyst/oil ratio and hydrocarbon residence time designed to particularly restrict the deposition of coke on the low coke producing zeolite catalyst, converting a second higher coke producing hydrocarbon 'feedin admixture with the suspension in the second conversion zone under operating conditions of temperature, catalyst/oil ratio and hydrocarbon residence time designed to particularly deposit coke upon the catalyst during the conversion operation, cyclonically separating the suspension at the outlet of each conversion zone into a hydrocarbon phase and a catalyst phase, recovering the separated hydrocarbon phase, and passing catalyst separated from the second conversion zone to a catalyst regeneration zone.
7. The method of claim 6 wherein the amount of regenerated catalyst combined with the catalyst separated from the first conversion zone is sufficient to provide a catalyst mixture having a temperature selected from within the range of 950F. up to about l,l00F.
8. The method of claim 6 wherein a greater catalyst to oil ratio is employed in the second conversion zone than employed in the first conversion zone.
9. The method of claim 6 wherein the catalyst mixture comprises less than 50 percent ZSM-5 type of conversion catalyst.
10. The method of claim 6 wherein a portion of the catalyst separated from the first conversion zone is combined with catalyst separated from the second conversion zone and thereafter passed to catalyst regeneration.
11. The method of claim'6 wherein a portion of the catalyst separated from the second conversion zone is recycled thereto and another separated portion is passed to catalyst regeneration.
12. The method of claim 6 wherein the first hydrocarbon feed is a low molecular weight fraction and said second hydrocarbon feed is a high molecular weight fraction.
13. The method of claim 6 wherein said first hydrocarbon feed is a virgin gas oil and said second hydrocarbon feed is selected from one or more hydrocarbon mature comprises from 40 to 80 weight percent of the ZSM-S type crystalline zeolite.