BE502216A - - Google Patents

Description

       

    <Desc / Clms Page number 1>
 



  IMPROVEMENTS IN THE REACTION OF REACTIONS WITH SOLIDS
FINELY DIVIDED, AND REFINING PETROLEUM OILS.



   The invention relates to the performance of reactions using finely divided solids in a fluidized state, and more particularly relates to systems in which a fluidized solid is employed in the processing, production or conversion of liquid. It is particularly applicable to systems, for example for the catalytic cracking of hydrocarbons, in which such a solid material.

   fluidized, for example a catalyst, is continuously returned in the cycle between a reaction zone or reaction vessel and a separate regeneration or reactivation zone or vessel., and especially where the atmosphere of the reaction zone is different of the regeneration zone and where there can be no mixing of atmospheres in the two zones.

   Such systems have been widely used commercially in the petroleum industry for the catalytic cracking of hydrocarbons.
Within the general problem of designing such a system, a number of different arrangements of receptacles and of displacement or transfer conduits have been proposed. When studying these designs, certain requirements must always be kept in mind. basic.

   The first of these is constituted by the fact that an adequate driving force must be foreseen for the circulation of the solid between the two zones or containers at the speed or in the desired proportion, with means for varying this driving force within the limits required for satisfactory control or control of operations.

   In addition, safe means must be provided to prevent back flow or return of the reagents from one of the vessels to the other through the transfer lines.
In the design of two-vessel systems of this type for use in the petroleum industry, especially in the catalytic conversion of hydrocarbons using a fluidized solid catalyst, the force of en =

  <Desc / Clms Page number 2>

 The drag commonly employed to circulate catalyst from one vessel to the other was the "fluistatic" overhead portion generated by a relatively dense column of aerated or gasified catalyst from a feed pipe.

   In a commercial installation model, these feed pipes are fed from each vessel of the two vessel system at or near the bottom of the bottom, and discharge through a variable orifice such as than a valve or gate valve. The rate or proportion of the catalyst output from the bottom of each of the supply pipes is controlled by an adjustment of the slide valve. The catalyst passing through these slide valves in each of the pipes falls into a low pressure zone where it is is removed by the stream of a suitable fluidizing agent, circulating at high speed, which agent is usually the reagent intended for the next phase of the cycle.

   This catalyst stream is then entrained as a dispersed or relatively low density slurry through a transfer line to the other vessel. The pressure drop across the slide valves is in each case considered sufficient to prevent the back flow of the fluidizing agent from the low pressure area into the supply pipe.
Various systems have also been suggested which do not have slide valves, and where internal conduits or feed pipes are provided for transporting fluidized solid directly from one contact zone to another, these two contact zones. being arranged separately within the same outer envelope.

   However, equipment of the type described in the above proposals has not been built or used commercially, despite the strong economic tendency to employ the simplest possible construction and operating controls.



   The reason for this is that a study of these models using internal supply pipes, and particularly where a spool valve is not provided, has shown a lack of flexibility in operation or presented a very range. narrow operating possibilities. A much more serious shortcoming, together with the need for a delicate balancing of pressures to keep these systems in operation, is the lack of adequate minimum guarantees against unintentional "temporary" reversal in operation, which could reverse the system. direction of flow in one of the interior supply pipes. In a catalytic cracking plant, this produces a mixture of hydrocarbon and air in the same vessel at concentrations well within the explosion limits.



  Even the slightest mixing of this type cannot be tolerated in a system where safety relies on a delicate balance of pressures, as the resulting sudden increase in temperature and pressure could easily aggravate the imbalance None of the simplified systems of these processes of the prior art does not have an adequate righting force to maintain the desired proportion or speed and direction of flow in the event of such a reversal, with a safety factor sufficient to compensate - check the pressure fluctuations which normally occur in the operation of such equipment.



   The low operational possibilities of these models are due to the fact that the normal pressure fluctuations in the operation of a fluidized solids treatment system are directly related to the linear velocity of the fluidizing agent o The type of calm flow and uniformity, which is essential for the satisfactory operation of such equipment, is achieved only with very low proportions of gas supply.



   On the contrary, feed ratios which are commercially useful for catalytic conversion processes usually result in the formation of a highly turbulent fluid layer within the contact zone characterized by temporary localized pressure fluctuations which constitute a large fraction of the total pressure drop across the layer.



   In any mass of fluidized solid, a gas bubble or a localized region of low density always tends to rise relative to the ma-

  <Desc / Clms Page number 3>

 tière entouranteo This trend extends everywhere across the continuum except where or. there is a definite obstacle or barrier of pressure drop sufficient to deflect the ascending bubble.

   An interior supply pipe, arranged vertically, not presenting an obstacle and open at its lower end, does not offer a positive barrier to the upward passage of such a bubble in the pipe, even if the density of the material confined inside the pipe is normally larger than that of the material outside A bubble entering the confined space at the lower end thereof will tend to rise unless the speed at which the fluidized solid passing downwards, out of the pipe, is clearly greater than the speed of rise of the bubble o The immediate effect of such a bubble inside the pipe is to decrease the differential density which constitutes the driving force or driving force of the circulation of the solid,

   thereby decreasing the rate at which the solid is delivered to the bottom of the feed pipe and increasing the tendency of the bubbles to pass upwards against the normal direction of flow. The same considerations apply to a single outer supply pipe, which is disposed vertically or at a slight angle to the vertical direction so as to provide direct downward supply from a receptacle into the vessel. other which is at a lower level, or when the second receptacle is placed below the first
In each case of such a direct feed pipe not presenting an obstacle, this tendency of the gas bubbles to interfere with the normal direction and the flow rate of the solid and the gas,

   is particularly important when the operations are carried out using a high linear velocity for the thinning agent, and where this agent is introduced into the layer of thinned solid at a level below the point at which such a pipe runs. feed opens in the diaper.

   Under these conditions, there is a high degree of turbulence inside the fluid layer, the number and size of gas bubbles or such low density regions inside the layer becomes very large, and a bubble whistically large to completely cover the cross section of the opening in the bottom of the supply pipe, is almost certain to be able to penetrate the pipe up to a certain point As such possibilities become more and more frequent with the increasing the gas velocity, there will be more and more bubbles or portions of bubbles, which will tend to rise in the high density zone ,. Thus, in a two-vessel system operating at any speed greater than the lowest fluidization speed,

   such an internal or external direct feed pipe can become an absolutely inefficient obstacle to the return mixing of the fluidizing gas from one zone to the other unless a slide valve is used in conjunction therein,
According to the present invention, finely divided solid is circulated between two containers without the use of a control slide valve.

   In this apparatus, each transfer line leading from one of the contact vessels of a two-vessel system into the other of the latter, is arranged in the form of a U-tube comprising a supply pipe. downward flow from the bottom of one of the receptacles, through a high pressure zone or safety bend which forms the bottom of the U-shaped part, and where the pressure is higher than it is everywhere elsewhere in the upper part of the system, and passing from there upwards into the other container, through an upward flow connector pipe As an added precaution, any reagent or fluidizing agent entering each contact zone through this ascending part of the U-tube connector,

   is introduced at a level which is clearly higher than the safety bend at the bottom of the U-shaped part connecting the two receptacles o This arrangement avoids return mixing of the type mentioned above without requiring the use of slide valves to create an obstacle of variable pressure drop, since both solid and gas are introduced upward into each contact zone, and there is no connecting line above the point of gas introduction, whereby gas bubbles could escape into the other container.

   The only passage for solids flow

  <Desc / Clms Page number 4>

 between the vessels is done by means of these U-shaped safety elbows which are placed at a level below the point where the reactive gas or the fluidizing agent is introduced into the connecting pipes.



   The normal direction of flow of the solids, in this arrangement, is maintained by establishing a pressure differential in each of the U-shaped connecting pipes, such that the total pressure exerted on the solid in the descending part exceeds the back pressure. total exerted on the solid in the part or ascending branch. This, in turn, is achieved by keeping the density of the fluidified solid in at least one of the two ascending branches significantly less than that of the solid in the descending branch of the same U, and by adjusting the top pressure in both re- containers so as to compensate for the differences in level of the catalyst layers located therein.

   The U-shaped safety bend at the bottom of each transfer pipe, below the point where the reagent or fluidizing agent is introduced, does not participate in the creation of the driving force for the circulation of the solid, and works essentially as a siphon tube with equal pressures on both arms of the safety elbow, at the same level, using minimum aeration to give maximum fluid density and maximum pressure safety effect., However, the length the entire ascending branch of the transfer line, including the ascending branch of the U-shaped safety section, acts as an ascending supply pipe to resist any reversal of the normal direction of flow of the solid and gas, thanks to the excess pressure existing at the bottom of this pipe, compared to the upper part.

   The value of the pressure differential between the top and the bottom of each of the U-shaped safety branches, as well as in the other parts of the system, is determined by the height and density of the fluidized solid contained therein. By this means, it is possible to mount the equipment taking into account the normal pressure fluctuations within either of the contact vessels. According to this invention, the pressure differential in these branches of safety, measured below the reagent injection point to the level of the bottom of the safety branch, is pushed to a value which is more than adequate to compensate for pressure fluctuations. normal, so as to give a safety factor for operations carried out with high gas velocities in the contact zone.



   According to another characteristic of the invention, a catalytic conversion process carried out in the manner described above forms one or more stages of a combined process for the conversion of crude oil into mineral oil products. of higher value, and particularly to achieve maximum yields of low boiling products in the boiling range of motor fuels and heating oils, and minimum yields of heavy products, such as fuel oil and tar, with a minimum of processing phases and significantly reduced investment and operating costs.

   This combined process comprises a series of integrated catalytic and / or thermal conversion and distillation stages, in which products from all phases are brought to a single fractionation stage and processed therein such that a product reduced crude obtained as a liquid residue of a crude product distillation phase, is washed with product vapors obtained in the crude distillation and conversion phases, and that particular streams containing all fractions desirable products, such as a particular stream of motor fuels, a particular stream of heating oils, and a particular stream containing all of the less desirable heavy products, such as fuel oil fractions,

   are recovered from said single product fractionation phase in optimum relative proportions, no subsequent fractionation of any of these streams being required for the recovery of the end products.



   In previously proposed combined distillation and crude oil conversion processes, the recovery with maximum yields of motor fuel products and heating oils, was usually achieved by subjecting fractions of light and heavy naphtha, derived from

  <Desc / Clms Page number 5>

 from the raw product retort to subsequent fractionation and if desired to appropriate thermal or catalytic refining treatments, such as reformation isomerization, hydroformation, alkylation, or by thermally or catalytically cracking a gas oil fraction from the product retort crude in order to subsequently recover a product with a low boiling point by subsequent fractionation of the products which have undergone the cracking,

   and subjecting the reduced crude product to further distillation at reduced pressures to produce additional tar and low boiling point products, primarily gas oil to be treated with the gas oil fraction from the retort. of raw product, as mentioned above. These processes require a series of fractionators to provide different product streams at desirable boiling ranges.

   For economical heat recovery, large numbers of heat exchanger devices are needed both in each unit and in combination between the units. Large vatting facilities must be provided to allow storage of the different products before proceeding to mixing in the desired proportions o The vacuum distillation equipment used to process the reduced crude product is costly from the point of view of investment, operation and maintenance o As a consequence of these complications, combined processes of the conventional type have to be employed on a relatively large scale in order to be economical. Normally, it is necessary to have refining capacities exceeding, say,

   approximately 20,000 bbl (barrel = 40 UoSoAo gallons) per day of crude product to make operations of this type profitable, while smaller refineries must be set up taking into account the often unavoidably high production of fuel oil and other heavy products of relatively low commercial value.



   According to this feature of the invention, the vacuum retort for the distillation of the reduced crude product is omitted, and significant savings in fractionation equipment as well as in ease of storage are achieved by bringing about a single fractionation phase substantially all of the vaporous products of the different phases of a combined process as well as the reduced crude product from the distillation of the crude product, and using these vapors to rid the crude product reduced to the fractionation phase of all its distillable constituents.



   For this purpose, the crude oil can be subjected to distillation in a conventional crude product distillation unit to produce an overhead stream of light virgin naphtha, a stream of heavy naphtha separated an even heavier stream of the kerosene and oil range. Diesel and Reduced Crude Product Bottoms Reduced crude product bottoms can still contain the fraction of gas oil intended, at least in part, to feed a gas oil cracking phase. Kerosene cannot be recovered as a product coming directly from the crude product retort since further conversion would not significantly improve its value as kerosene or diesel oil.

   All other fractions are supplied to a single fractionator substantially as follows
The reduced crude product obtained as described above is sent directly to an upper part of the lower contact section of a substantially conventional fractionation column. The heavy naphtha stream is subjected to catalytic or thermal reform at high temperature. rature or other conversion resulting in an improvement in its engine fuel qualities, and the total vaporized effluent from this conversion stage is fed to the fractionation column at a point below the feed point of the engine. crude product reduced substantially at the temperature of the conversion phase.

   The light virgin naphtha fraction is likewise supplied to the fractionation column at a point below the feed point of the reduced crude product, and can be injected either in a liquid state or in a vapor state, after appropriate reheating, if desired Different end product streams are recovered from the fractionator, for example a fuel gas head, a low boiling fraction of the boiling range of engine fuels, a fraction of heating oil, a fraction of gas oil and a fraction of heavy tailings from the range of fuel oils o In addition,

   he can be pro =

  <Desc / Clms Page number 6>

 produces a gas oil fraction which contains the virgin gas oil fraction desired for feed to other conversion equipment, together and in admixture with a recycle or converted fraction, in an approximately optimum ratio, for a final conversion operation.

   Most of the gas oil fraction is sent to a preferably catalytic cracking phase to be converted there into additional amounts of engine fuel, diesel oil, gas oil range cyclisation feed and of heavy tail products. The total effluent from this cracking phase is likewise sent without appreciable loss of. heat to the fractionator at a point below the reduced crude product feed, but preferably above the reformation product feed points from processing heavy naphtha and light virgin naphtha. In this way, the reduced crude product or its equivalent is successively subjected to countercurrent heating, to vaporization and to the scrubbing or washing action,

   first with vapors from the gas oil cracking process, which are heavier and may be at a lower temperature than the naphtha vapors used, and then with low molecular weight naphtha vapors and preferably at higher temperature o
By operating substantially as described above, extremely large volumes of vapors from the process are available and used to purify the reduced crude product with the effect of keeping the volume of heavy fuel ultimately produced to a minimum. as low as and,;, if;, 'Desired even lower than that which may be obtained in conventional operation involving vacuum distillation of the reduced crude product.

   Total product naphtha is prepared as an overhead product which has been maintained at elevated temperatures for a minimum time and is, therefore, of superior quality in machine cleanliness. A single, heavy fuel fraction is obtained in which all the heavy constituents formed in the different phases are combined, and which can be subjected to a single filtering treatment to recover a final fuel oil of excellent quality. These advantages are to be added to the savings evident in the heat exchanger equipment, resulting from the elimination of the various intermediate heating and cooling operations,

   and the savings in intermediate storage resulting from the use of a single phase of product fractionation.
The number and type of intermediate conversion stages employed in accordance with this feature of the invention will depend entirely on the character of the crude oil to be treated and on the titers, quality and relative proportions of the desired products. can be added or substituted for those which have been mentioned. For example, a coking or viscosity reduction treatment can be applied, either to the reduced crude product before entering the fractionator, or to fuel oil recovered from the fractionator, depending on process requirements and required products.

   The first process may be employed when it comes to increasing the feed supply for catalytic cracking. The other process may be dictated by the specific fuel oil requirements of the product, which may be affected. by the constituents formed in the crackingo phase Other conversion phases which may be incorporated into the process of the invention include hydroformation and / or other methods of improving the octane number. for virgin naphtha, rupture of thermal or suspenscid viscosity of reduced crude products, etc ...



  Following the basic principle of this feature of the invention, the product vapors of all of these phases can be brought to the single fractionation phase substantially as described above.



   In some cases the process can be further simplified by removing the retort and feeding all the crude oil directly into the fractionator. This is possible if it is not necessary to have separate fractions of virgin naphtha and / or kerosene. By operating in this fashion, all of the raw product can be fed to an upper portion of the lower contact section of the product fractionator. Heavy naphtha stream and gas oil cracking charge can be removed from the fractionator and

  <Desc / Clms Page number 7>

 subjected to the conversions described above, particularly to reforming and catalytic cracking.

   The vaporized products from these conversions can then be returned to the fractionator at a point below the raw product feed point. Light naphtha and overhead gases from the fractionator can also be returned to the bottom of the fractionator. - tor to help wash the raw product.

   In all other respects, the operation of the process will be substantially analogous to what has been indicated above ,,
By carrying out the combined refining process described above, the production and quality of the obtained engine fuel fraction suffers to some extent due to thermal regeneration or catalytic reforming in the absence of large amounts of hydrogen. should not be used. rather than catalytic hydroformation at the reforming stage to allow efficient operation of the fractionator.

   Catalytic hydroformation or similar processes are far superior to reforming processes carried out in the absence of hydrogen in product quality, including production-to-octane ratio, sulfur content and grades. cleanliness of the gasolineo However, processes of this type result in the net production and recycling of large quantities of hydrogen. The introduction of this hydrogen with the reforming product into the common fractionator seriously hinders the recovery of the value light hydrocarbon bottoms or necessitates an uneconomical size fractionator model.



   According to a variant of this characteristic of the invention, the combined refining process described above can comprise, as one of the intermediate phases, the step of catalytic reforming of the naphtha fractions obtained in the various phases in the presence of hydrogens. and also the use of at least a part of the hydrogen produced during this reforming step to treat other suitable fractions;

  obtained in the various phases of the combined process outside said reforming phase and from the purifier or "stripper" of the fractionator. The preferred embodiment of this variant of the invention comprises the incorporation of poration to said simplified combined process of catalytic naphtha reforming operations of the type known as platinum reforming (Platforming) in which the virgin and / or cracked naphtha, coming from the different phases of the Combined process is contacted with a platinum catalyst at elevated temperatures of about 800-1,000 F,

   high pressures of about 400-1,000 pounds per square inch and relatively high hydrogen recycle proportions of about 3,000 to 12,000 standard cubic feet per barrel maintained by recycling the hydrogen produced in the process . Under the specified conditions, three main reactions take place, namely the dehydrogenation of the naphthenes to the corresponding aromatics,

     hydrocracking of high molecular weight hydrocarbons to saturated low molecular weight hydrocarbons and isomerization of naphtha and straight chain hydrocarbons o In addition, dehydrocyclization of straight chain hydrocarbons directly into aromatics and desulfurization takes place substantially all of the sulfur to hydrogen sulphureo Mainly as a result of the high partial pressure of hydrogen, the catalyst retains its activity for many months without regeneration.



   The combined process can include other conventional hydroformation processes. -For example, virgin and / or cracked naphtha can be brought into contact in the presence of external hydrogen and / or hydrogen formed in situ, with well-known catalysts such as molybdenum oxides. , aluminum, vanadium-and tungsten, with or without silica, supported on alumina or cobalt moly'odate as it is, or molyb = dine supported on zinc aluminate, etc. ...., at temperatures of about 800 to 10,000 F, pressures of about 50 to 10,000 pounds.

   square inch and recycle gas ratios of about 10,000 to 3,000 standard cubic feet / barrel of a recycle gas containing about 80% H2. Continuous or intermittent regeneration of the catalyst can be used, if necessary, to remove carbonaceous deposits by counter-burning.

  <Desc / Clms Page number 8>

 produced with air mixed with recycled flue gas from the regeneration process.

   Regeneration temperatures are generally limited to 1,000 - 1,200 F at any suitable pressure. Platinum forming treatment or other conventional hydroformation process can be carried out in a fixed bed type operation. with a moving bed, of the fluid or suspensdides type, all in a manner known per se.



   According to a further embodiment of the invention, when such catalytic hydroformation processes are incorporated into a combined process of the type specified above, the hydrogen made available in the hydroformation phase is employed to improve the cracked fractions removed from the combined stripper-stripper by an improvement treatment including hydrogenation o For example any desired portion of the gas oil recycle feed can be treated either with separated hydrogen from reformer's effluent,

   either with the total effluent from the reformer o The treatment with hydrogen can be carried out under non-cracking conditions comprising temperatures of about 500 to 900 F and pressures of about 750 to 4,000 livo / square inch in the presence of conventional hydrogenation catalysts, such as heavy metal oxides or sulfides of Groups V, VI and VIII. Such hydrogenation improves the quality of the recycle feed such as catalytic cracking feed by hydrogenating coke-forming constituents, such as polycyclic aromatics, prior to cracking.



   The hydrogen coming from the hydroformation phase can also be used to improve the light gas oil which has undergone cracking in the boiling ranges of heating or diesel oils. For this purpose, oil of the type mentioned can be used. hydrogenated on the catalysts of Groups V, VI and VIII just mentioned, either substantially under conditions of desulfurization, such as temperatures of about 500 to 800 F and pressures of about 100 to 500 liv.

   square inch, or substantially under hydrogenation conditions similar to those described with reference to the hydrogenation of the gas oil recycle feedstock Rather than employing separated hydrogen, the total hydroformation effluent may be used to improve light gas oil which has undergone cracking, as will appear more clearly below
The reformed naphtha can be mixed with the naphtha fractions recovered from the stripper-fractionator,

   preferably after these have been subjected to conventional finishing treatments such as caustic washing and / or suitable scrubbing methods known in the art. The reformed naphtha itself is of such quality and purity. that it does not require further finishing of this nature
Although these types of operations may require special means of fractionation and / or gas-liquid separation in addition to the combined fractionator-stripper, the resulting increase in investment cost is at least partially compensated for. by the improved quality of the final engine fuel mixture. In addition, the naphtha finishing equipment can be considerably reduced in size, which ensures an appreciable saving in the capital cost.



   According to another embodiment of this variant of the invention, the total effluent of the catalytic naphtha hydroformation phase, comprising all of the hydrogen present in this phase, is used without intermediate cooling to purify at least a part of the reduced crude product, having its entry into the combined stripper-stripper, in a separate purification phase treated substantially at the pressure of the hydroformation phase.

   Assuming a temperature of the reduced crude product recovered from the retort, of about 600 to 750 F and a temperature of about 800 to 1000 F for the hydroformer effluent, this preliminary scrub can take place in temperature limits of about 650 to 800 F at a pressure of about 100 to 1,000 livo / square inch. Under these conditions, a significant proportion of the gas oil and lighter constituents of the reduced crude product, about 5 to 50%, will go overhead with the hydroformate and hydrogen. The hydrogen can then be separated from the product. normal head

  <Desc / Clms Page number 9>

 liquid and returned to the cycle to the hydroforming phase after appropriate reheating.

   The hydroformate containing the stripped elements of the reduced crude product can be sent to the fractionation section of the combined stripper-stripper or be fractionated in a separate tower to produce gas oil, for use as a catalytic cracking feed, and hydroformate which can be mixed with the finished naphtha from the combined stripper-fractionator as described above.

   The reduced crude product which has been pre-purified is sent to the purification section of the combined stripper-stripper for further purification with the effluent vapors from other phases of the process.
In all of the embodiments described above, excess hydrogen made available at the hydroformation phase, after satisfying the recycling requirements of this phase, can be sent to the scrub section of the stripper-fractionator. combined to aid in the purification of the reduced crude product The hydroforming naphtha can be either virgin heavy naphtha alone or heavy naphtha extracted from the combined stripper-stripper and containing both virgin and cracked naphtha ,

   or any suitable mixture of such heavy naphtha. It may be desirable to make all virgin naphtha available for stripping the reduced crude product in the combined stripper-fractionator. In this case, a combined stream of heavy naphtha and naphtha virgin light can be sent directly to the scrub section of the stripper-fractionator, and only a fraction of heavy naphtha from the stripper-fractionator be subjected to hydro = catalytic formation as described above;

   
The general nature and the various subjects of the invention having been explained, it will be better understood in the course of the detailed description given below with reference to the accompanying drawings, in which
FIG. 1 is a schematic representation of one form of apparatus intended for carrying out the treatment, and comprising a reactor and a regenerator suitable for the catalytic cracking of hydrocarbons.
Figure 2 is a variant of a model comprising the same fundamental principles
Figures 3 and 3a show a schematic flow plan of the combined process according to the invention.



   Figures 4 and 4a show a modified schematic flow plan of the variant of the combined process including the hydroforming phases of naphtha.



   Figure 5 shows an alternative hydroformation phase given in more detail.



   Figure 6 shows a variant of the arrangement shown in Figures 3, 3a, 4 and 4a.



   In Figures 3 to 6, the same reference notations designate similar parts.



   Referring now to Figure 1 of the drawings, it can be seen that the illustrated system comprises a reaction vessel or cylindrical reactor 10, arranged vertically, containing a layer or fluidized bed 11 of catalyst, the level of which is indicated at 12, in which the hydrocarbons introduced through line 13 are subjected to catalytic cracking. Hydrocarbon gases or vapors passing upwardly through fluid bed 11 maintain the bed in a dense turbulent fluid state giving it the appearance of a boiling liquid.

   A separate regeneration vessel or regenerator 14, disposed at approximately the same level as reactor 10, contains a similar fluidized layer 15 of catalyst, having a level indicated at 16, which is subjected to regeneration to ignite deposits. carbonaceous substances formed during the cracking reaction. The regeneration gas is an oxidizing gas, preferably air which passes upwardly through bed 11 so as to maintain it in a dense turbulent fluidized state.

  <Desc / Clms Page number 10>

 



   The catalyst used for this process can be a silica-based catalyst, prepared by the acid activation of bentonite clays, or a synthetic catalyst derived from silica gel (silica gel) or other forms of silicic acid. . The catalyst may be of the silica-alumina or silica-magnesia type, with suitable additions of other active constituents such as zirconia, borie or the like. This catalyst may be in the form of a finely divided powder prepared by grinding, or in the form of small spheres prepared by suitable drying methods, in the case of synthetic catalysts. The catalyst preferably contains particles having particle size limits including particles in the size range of 0 to about 200 microns in diameter.

   The top pressure in the reactor can be about 6 to 10 lb./sq. Inch as determined by the back pressure of the product recovery system (not shown), i.e. the drop in pressure. pressure which the product vapors must overcome The pressure in the regenerator can be about 6 to 15 pounds, and the ratio of the reactor pressure to the regenerator pressure is an important operational control, as one set out in more detail below. Both the reactor pressure and the regenerator pressure can be increased considerably beyond the values given here, maintaining the proper ratio between them and can be 100 lb. square inch or higher.



   The feedstock for the catalytic cracking process can be gas oil, naphtha, heavy distillate, overhead crude, whole crude, or other crude oil fractions separately or in combination, or the process can be applied to liquid hydrocarbons or to mixtures of hydrocarbons derived in part from sources other than petroleum.



   The temperature of the cracking reaction can be between 700 and 1,100 F, preferably 900 to 950 F, and the regeneration temperature can be between around 900 to 1,200 F, preferably around 10,100 to 1,150 F. The system is suitable to operate under thermal equilibrium conditions, that is to say that the combustion of carbonaceous deposits formed on the surface of the catalyst during the cracking reaction, gives sufficient heat for the regenerator and the reactor to supply this produced heat to allow vaporization and cracking to take place at the desired temperature level.

   The catalyst-to-oil ratio required to maintain this thermal equilibrium operation will vary with the characteristics of the feedstock, the temperature to which this feedstock is preheated by indirect heat exchange with different currents. effluent from the cracking and regeneration processes, and the desired cracking temperature. This catalyst-to-petroleum ratio can vary from about 5 to 1 to 30 to 1 parts by weight, and can, under preferred conditions, be about 10 to 1.



   Spent catalyst is removed from reactor 10 and flows downward as a dispersed suspension through stripper 17 which may be provided with baffles or disc baffles 18. Line 19 supplies scrubbing vapor or another scrubber gas at the bottom of the stripper. The catalyst removed from the bottom of the stripper 17 passes through a U-shaped safety branch 20 to enter the connector riser pipe 22. The riser feed pipe 22 extends into the bottom of the regenerator. 14 and ends in the upper part of the fluid bed 15 which is contained therein, below the level 16 of this bed.

   Aeration intakes 21 are provided in this U-shaped safety branch 20 which connects the stripper 17 to the ascending supply pipe 22, to supply the minimum quantity of steam or other aeration gas required to maintain fluidity. in the safety branch with maximum catalyst density. An inverted deflector cone 23 may be provided in bed 15 a short distance above the top of upstream feed pipe 22 to provide better distribution of catalyst entering bed 15.



   Catalyst level 16, inside the regenerator, is secured by the open upper end of a discharge pipe. Descending 24,

  <Desc / Clms Page number 11>

 which extends upwardly into the regenerator 14. The catalyst from the fluid bed 15 overflows into the upper part 25 of this pipe 24, which has a larger diameter than the lower part of said pipe 24 to serve as a well. discharge and provide backflow capability to compensate for small flow fluctuations in the catalyst flow. The total amount of catalyst maintained in the fluid bed is thus substantially constant for a given catalyst, allowing for slight variations resulting from changes in the density or particle size distribution characteristics of the catalyst.

   Vertical slots 26 may be provided at the top of the discharge well 25 to give a more uniform flow rate to the catalyst discharge, and to allow slight variations in the level 16 of the catalyst without causing large fluctuations in the following flow rate. which flows the catalyst into the discharge well.



   The catalyst flowing down through pipe 24 and passing through U-shaped safety branch 27 is maintained at maximum fluid density, with the aid of the minimum amount of steam or other gas. aeration introduced by the aeration intakes 28. The safety branch 27, which functions to prevent any tendency of the hydrocarbons injected through the line 13 to flow backwards through the line 27 into the regenerator, comprises the part of the conduit, connecting the regenerator 14 to the reactor 10, which is located below the level of the oil injection conduit 13, comprising all the curved portion of the bottom of this conduit located between the conduit 13 in the riser and a corresponding arbitrary point 29 of the down pipe 24.

   Point or level 29 is found. between the straight vertical descending portion 24 of this conduit and the curved safety branch portion 27, at approximately the same level as the point where the oil injection conduit 13 enters the ascending transfer conduit 30.



  The catalyst flowing downward through pipe 24 and safety branch 27 then passes into reactor 10, through upward transfer pipe 30 which is of a somewhat larger diameter than it. of the pipe 27 and is joined to the pipe 27 at the level of the oil injection pipe 13 above the U which forms the safety branch.



   Gas oil or other hydrocarbon feedstock for the cracking process, which can be preheated by indirect heat exchange with different product streams or regeneration aeration gas in any suitable manner, is injected. in line 30 via line 13 at a point located just above the level where line 30 meets the U-shaped safety branch 27. The mixture of oil and catalyst formed by the injection of liquid hydrocarbon preheated through line 13 passes, as a dilute dispersed suspension of cracking catalyst in oil vapors, through line 30 into reactor 10. The very large increase in gas volume caused by vaporization and expansion oil in this line, gives the resulting suspension a high speed and low density.

   The high velocity of the catalyst-oil slurry in line 30 decreases sharply with increasing vessel diameter where this transfer line opens into reactor 10. Thus, the density in transfer line 30 may be reduced. have a value within the limits of 2 to 10 pounds / cubic foot under conditions where the density of catalyst in bed 11 and stripper 17 may be about 20 to 30 pounds / cubic foot and the density in pipe 24 and safety legs 20 and 27 be about 35 to 45 pounds / cubic foot.



   The transfer line 30 passes through the side wall of the reactor 10, near the top of the scrub section 17, and terminates in the reactor in an inverted distribution cone 31. The dilute suspension of catalyst in oil, which is located in the pipe 30, is introduced into the bottom of the fluid bed 11 through a perforated grid 32 placed horizontally at the upper part of the distribution cone 31.

   An inverted deflector cone 33 may be provided within distribution cone 31 and below grate 32 to provide better distribution of catalyst and hydrocarbon vapors through the cross section of the bottom of the tank. 11 fluid bed upper end of the distribution cone

  <Desc / Clms Page number 12>

 31 has a smaller diameter than the inside diameter of reactor 10 to leave an annular space 34 through which catalyst can flow downward from reactor 10 into the top of stripper 17.



   In this arrangement, the dense catalyst from regenerator 14 into downpipe 24, including discharge well 25, acts in the usual manner to create fluidistatic pressure in the pipe, and the direction of flow of catalyst from. regenerator 14 to go to reactor 10 is maintained by the equilibrium of the fluidistic pressure in the safety branch 27. The fluidistatic pressure exerted by the dense catalyst in the pipe 24 does not exceed that exerted by the dilute suspension of catalyst in the upward transfer line 30 only by the small pressure drop due to frictional losses in the safety branch amounting to a fraction of a pound per square inch.

   This small pressure drop in the safety branch can hardly be greater, in an open tube, in the absence of an arbitrary pressure drop obstacle such as the slide valves commonly employed in other systems to control the pressure. catalyst flow rate. Under these circumstances, any additional catalyst falling into the discharge tube 25 displaces solid in the transfer line 30, causing a more intense rate of the catalyst circulation.

   Likewise, the density of the dense suspension of catalyst in the riser 22 leading to the regenerator 14 is variable within fairly wide limits which may be of the order of 10 to 25 pounds per cubic foot, but in any case. is appreciably less than the density in the reactor and the stripper, so that the fluidistatic pressure exerted on the downward stream of catalyst, at the bottom of the stripper, tends to exceed the fluidistatic pressure exerted on the upward catalyst, at the bottom of the pipe ascending 22.



   The density of the catalyst in the riser pipe 22 is regulated by injecting air through a fan 35 or by injecting water vapor or another gas at a controlled rate, through a pipe 36 opening into the bottom of the pipe. pipe 22, just above the level where pipe 22 is connected to safety branch 20.

   An increase in the air flow rate supplied by the fan 35 therefore results in an increase in the proportion of catalyst circulating between the vessels, due to a decrease in the density of the catalyst column in the riser 22 and a decrease in the fluidistatic return pressure exerted by the latter, When air is used as the gas flow controlling the flow rate, and introduced through the pipe 36, the fan 35 can take its suction from the main air duct 37 supplying air to regenerator 14 through a fan 38.



   The driving force of the circulation of the catalyst between the vessels arises, in this system, from the fluidistatic pressure equilibrium as indicated above, in that the density of the catalyst mixture in the down pipe 24 substantially exceeds that. prevailing in the transfer line 30, and that the density of the catalyst mixture in the stripper 17 and the reactor 10 substantially exceeds that prevailing in the riser 22.

   Looking more closely at this equilibrium, it can be seen that the total pressure exerted on the reactor side of the safety branch 20, which results from the gas pressure at the top of the reactor 10 plus the fluidistatic pressure of the catalyst in the bed 11 and the stripper 17, can be used to exceed the back pressure exerted at the base of the riser pipe 22 on the regenerator side of the safety branch 20, which results from the pressure of the catalyst in the riser pipe 22 plus the small pressure due to the portion of the catalyst in bed 15 above the point where line 22 enters regenerator 14, and the back pressure of gas above bed 15 to regenerator, causing a flow of catalyst which tends to equalize the pressure differential thus created.

   On the other hand, the gas pressure at the top of the regenerator plus the catalyst pressure inside the downpipe 24, exerted on the regenerator side of the U-shaped safety branch 27, varying with the level of the catalyst in the tube. the evacuation shaft 25 and increasing

  <Desc / Clms Page number 13>

 with the catalyst overflow into the discharge well, counterbalances the back pressure of the oil-catalyst slurry in transfer line 30 plus the fluidistatic pressure of the catalyst in bed 11 and the back pressure at the top of the reactor, and exceed it only by a small pressure difference equal to the pressure drop due to friction.



   The U-shaped safety branches 20 and 27 essentially constitute the part of the transfer pipes coming from the reactor and from the regenerator which is located below the level of the gas injection pipe 36 and of the oil injection pipe 13 respectively, Thus, the safety branch 20 comprises the entire U-shaped pipe between the pipe 36 for the introduction of the gas stream controlling the flow, at the base of the ascending pipe 22, and the corresponding arbitrary level 39, located in the downward flow of catalyst, below reactor 10, at a point near the bottom of stripper 17. The fluidistatic pressure at these two points tends to be equaled or offset by the fluid behavior of the catalyst slurry therebetween.

   The ascending portion of this safety branch, between the bottom of the U 20 and the air injection pipe 36, acts as an obstacle or positive pressure safety opposing any forces that could tend to flow back the air coming from the riser pipe 22 towards the reactor 10, against the normal direction of flow. This obstruction or safety effect exists because the fluidistatic pressure at the base of the U-shaped part is greater than it is at the point of injection of air at 36.



  The descending part of the safety branch 20, between level 39 and the bottom of the U-shaped part, creates a pressure head sufficient to counterbalance the pressure differential which produces the effect of safety or obstruction in the valve. rising part of the safety branch 20, so that the net effect of the safety branch on the circulation of the catalyst is that of a siphon tube The situation of level 39 inside the stripper 17 is slightly higher to that of the level of the air injection line 36, as a result of the drop in frictional pressure of the stream of fluidized solid flowing in the safety branch 20, and because any catalyst in the bottom part of the stripper 17, which may be part of the descending branch of the U, has a fluid density somewhat less than the maximum.



   Similarly, in the safety branch 27, which extends from approximately the level of the oil injection line 13 to the corresponding arbitrary point or level 29 in the down pipe 24, all. tendency of the oil vapors, in the ascending transfer line 30 to flow in the opposite direction, towards the regenerator 14, meets the opposition of the fluidistatic pressure head of the catalyst contained in the ascending part of this safety branch , between the level of the injection pipe 13 and the base of the Uo part. Thus, there is no possibility either for the main stream of regeneration gas injected through the pipe 37, or for the fluidization stream injected into line 36,

   to appear as gas bubbles at the base of the descending branch of the descending U-shaped portion of the reactor side of the system, without first overcoming the positive pressure obstruction established in the ascending branch of that portion The same considerations apply to the oil vapors in the transfer line 30 vis-à-vis the U-shaped portion 27 and down pipe 24 of the regenerator.



   The fluidistatic pressure head of the catalyst contained in these U-shaped parts therefore serves as an obstacle or positive safety towards the return mixture of the oil going from the pipe 13 to the regenerator 14 or of the air going from the pipe. 36 in the reactor 10. This effect of safety or obstruction of these U-shaped parts, below the entry point of the oil injection pipe 13 and of the gas pipe controlling the flow 36, provides a safety factor with respect to and in addition to the pressure head component and the speed of the catalyst flowing downward in each of the safety branches.

   These latter factors are in themselves sufficient to maintain the flow of the catalyst in the desired direction in the absence of pressure fluctuations in the system, but are not sufficient to

  <Desc / Clms Page number 14>

 to prevent with certainty a reversal of this direction of flow as a result of normal pressure back-strains within the system. As indicated above, the depth of the safety branches required for this purpose should be sufficient to compensate for the pressure fluctuations which occur in normal operation.

   In conventional catalytic cracking plants, where slide valves are used to control the proportion of catalyst circulation, pressure backpressures in the order of 2 to 3 pounds per square inch are encountered, particularly in risers of the dilute phase where the catalyst falling through the slide valve is taken up by the thinning agent for the next phase of the operating cycle and transported upwards to the next contact vessel as a slurry dispersed diluteda In the operation of such a system, the pressure drop across the gate valve is never less than about 3 pounds per square inch.

   According to the present invention, slide valves are not used to control or adjust the proportion of catalyst circulating, and the pressure drop which is required to prevent back-mixing of gases from one of the contact vessels to. the other is supplied by the U-shaped safety branch below the level of introduction of the two fluidizing gases. So, for example, a fluidized catalyst having a density of 38 pounds per cubic foot under minimum aeration conditions, will require a pipe height of approximately 3.8 feet per pound of pressure drop, and a height of 15 feet. in these U-shaped safety legs will give a safety pressure differential of about 4 pounds per square inch.

   This can be compared to a total pressure drop of about 2 to 8 pounds per square inch across fluid beds 11 and 15 of the reactor and regenerator, depending on the depth of these fluid beds and the density of the catalyst slurry. that is in those.



   A further feature of this invention is that the oil feed stream injected into line 13 enters transfer line 30 with the same direction of flow as the catalyst therein. finds, so that there is no abrupt change in the direction of flow of the catalyst, either at the point of injection of the oil feed or at the point where the catalyst stream enters this transfer line . This concurrent flow of oil from line 13 and catalyst from line 27 through transfer line 30 into the reactor provides smoother operation with much smaller pressure fluctuations at this point than those observed in the reactor. the operation of equipment controlled by conventional slide valves.

   Similarly, in line 36 at the base of riser 22, the flow of catalyst and the flow of introduced gas are collinear, so as to minimize the pressure fluctuations which tend to occur as a result. the formation of poorly dispersed catalyst sludge in the transfer line where there is a sudden change in the direction of solid flow at the point of gas injection. Pressure fluctuations in pipe 22 are also minimized by maintaining the catalyst column therein as a dense slurry rather than a dilute slurry.

   Any weakly dispersed catalyst sludge persisting for a short distance during the upward path of the solid within this pipe will therefore only cause small local pressure backstrokes as a result of the difference in density, instead of Marked back pressure surges that can occur when such sludge occurs in a dilute phase transfer line.



   The main supply of combustion air for the regeneration process is introduced into the bottom of the regenerator 14 through line 37 coming from a fan 38 and through an auxiliary burner 40. The auxiliary burner 40, arranged at the base of the regenerator 14, is used to supply heat to the catalyst during the start-up period. For this purpose, line 37 opens directly into burner 40, and a separate supply of gaseous or liquid fuel, not shown, is provided during start-up.

   During normal operations, there is no combustion in this burner, A hollow hemispherical refractory cap or mask 41 covering the

  <Desc / Clms Page number 15>

 upper part of the outlet of the burner 40 protects the metal gas distribution grid 42, supporting the fluid bed 15, against direct radiant heat when the burner is in use. The tight spacing 43 between cap 41 and the protruding top of burner 40 provides a purge device for the vapors and also gives a high velocity to the gas stream at this point, which helps to prevent catalyst from being trapped. the cone at the bottom of the regenerator, below the grid 42, to fall into the burner.



   The design of the grid 42, supporting the fluid bed 15, is preferably such that there is a sufficient pressure drop across this grid to prevent backflow of catalyst through the openings therein. whereby the cone at the base of the regenerator 14, below the grille 42, results in usually working without a catalyst when the fan 38 is providing a normal air supply. The riser pipe 22 is shown as protruding into the fluid bed 15 to a level situated above the grid 42, which adds: the difference in static fluid pressure of the catalyst included in this part of the pipe, compared to that prevailing in the fluid bed 15, to the driving force of the circulation of the catalyst, and thus decreases the total height of the required container.

   In other cases, however, it may be desirable that the pipe 22 terminate at the level of the grill 42 or at an even lower level, such as at the point where it enters the bottom of the container. regeneration 14.



   The top pressure inside the reactor 10, above the fluid bed 11 contained therein, may be about 9 pounds per square inch, as indicated above, this being determined by the prevailing back pressure. in the product recovery system. The product vapors which pass over from the fluid bed 11, contain particles of entrained catalyst which are separated from the product vapors by two or three phases of a suitable gas-solid separator 44, such as a cy separator. - clone, a "multiclonic" device or the like, and returned from it to bed 11 through immersion branches. The vapors are then transported through line 45 to the product recovery system.

   If desired, small amounts of residual entrained catalyst, collected in the recovery system as catalyst slurry in oil, can be returned to the cracking system through the injection line. oil 13 or any other convenient point.



   The top pressure exerted on the regenerator 14, above the fluid bed 15 contained therein, can be about 11 pounds per square inch and be controlled by a control valve 46 disposed in a chimney 47 of the regenerator. .



  Exhaust gases from the regeneration pass into the chimney 47 via the gas-solid separator device 48 such as a cyclone separator. a "multiclonic" apparatus or the like, which may be arranged as a two-phase or three-phase separation system with dip legs to return the catalyst thus recovered to the fluid bed 15.



   The adjustment of valve 46 may be controlled by a differential pressure control 49 suitably connected to the top of regenerator 14 and reactor 10 so as to maintain the desired pressure differential between the tops of these vessels. The pressure differential between these vessels, which is part of the pressure balance in the system, is thus maintained at a constant value.

   Although the present example is described in connection with a pressure at the top of the regenerator somewhat higher than the pressure at the top of the reactor, in order to achieve certain advantages incident to a regeneration at a higher pressure level, exactly the same principles can be achieved. apply to an operation in which the two vessels are at the same pressure at the top, or in which there are differences between them in either direction.



  The invention also contemplates operating with one of these vessels disposed at a higher level than the other, with appropriate adjustments of the pressures at the top and the length of the two pipes going to and from the higher vessel.

  <Desc / Clms Page number 16>

 
The main operating variables on which the control of the operation of this system is usually based include the supply of air to the riser 22, the total quantity and quality of the catalyst in the system, and the temperature as well as the quantity of oil supplied to the injection line 13.

   To take full advantage of the available regeneration capacity, it is most economical to run the main air fan 38 at constant speed, in capacity, and to make any adjustments that may be required in the section of the fan. reactor to produce the amount of coke on the catalyst that will keep the system in thermal equilibrium at the desired cracking temperature and with maximum coke burning capacity. It is possible, however, to make certain adjustments to the rate at which the main combustion air is supplied by the blower 38, if such adjustment is warranted in special circumstances.



   Although the total amount of catalyst maintained in the fluid bed 15, within the regenerator 14, is substantially constant as a result of the fixed level 16 at the top of the discharge well 25, the catalyst contained in the discharge well. The discharge finds a free level 50 somewhat below level 16 of bed 15, depending on the proportion of catalyst circulating and the pressure differentials established elsewhere in the system. The existence of such a free level allows considerable adjustments to the circulating volume without causing the system pressure imbalance, and the difference between level 50 and level 16 is a measure of the driving force. reserve for the circulation of the catalyst, which allows such adjustments.



   The catalyst in the reactor is directly controlled by the total volume of catalyst in the system. Changes in this volume of catalyst can be made to produce a change in the level of fluid bed 11 in reactor 10 by adding makeup catalyst or removing spent catalyst at any convenient point, for example in the pipe. 22 leading to the regenerator. Some of the catalyst removed in this manner can be vented to allow fresh catalyst to be added to maintain the desired level of activity in the total circulation stream, or it can be entirely stored for reuse.



   The temperature in the reactor 10 is normally controlled by controlling the flow rate of the catalyst circulating in the system. In order to accomplish this operation, the candence at which the air is supplied to the pipe 22 by the blower 35 is controlled by a regulator. temperature indicator 51 disposed at a suitable point in bed 11. In the event that the reactor temperature thus determined tends to fall below the desired value, this mechanism is arranged to increase the air supply. to pipe 22, by reducing the density of the catalyst slurry therein and by increasing the flow rate of the circulating catalyst.

   For a higher flow rate of circulating catalyst, there is more hot catalyst flowing per unit time in well 25, coming from regenerator 14, and this catalyst passes through reactor 10, increasing the temperature. of the reactor and the severity of the cracking for a given oil feed rate so as to produce an additional amount of coke on the catalyst and a higher proportion of heat production when this coke is removed by combustion of the catalyst in the reactor. the regenerator 14.

   On the other hand, too high a temperature in the reactor will act by the same mechanism to cause a decrease in the supply of air to pipe 22, resulting in a decrease in the proportion of circulating catalyst, a less supply of heat to the pipe. reactor and lower coke production per unit time. The amount of oil supplied through line 13 and the temperature to which this feedstock is preheated prior to its injection into the catalyst stream can also be controlled by controlling the temperature of the reactor. This preheating of the oil can vary from about atmospheric temperature to 750 F or more, and is usually on the order of about 400 to 650 F.

   The severity of the cracking, which determines the distribution of the product and the quality of the product, is a factor in all these variables,

  <Desc / Clms Page number 17>

 comprising the feed rate, type and temperature of the oil feedstock, and the circulation rate, activity and temperature of the catalyst stream entering the reactor
As a safety feature of this operation, the flow of catalyst from the ascending branch of the U 27 into the transfer line 30 is measured by a pressure drop monitor 52. mounted with a venturi 53.

   A safety valve 54, which operates in the fully open state under normal conditions, can also be arranged in line 27 at this point. There is no throttling action caused by this. valve which operates in either fully open or fully closed condition, and a venturi throttle 53 is constructed so as to create only the smallest pressure drop that can be reliably measured,

   pressure drop on the order of a few tenths of a livery The controller 52 is connected to the valve 54 so as to shut off the flow of catalyst completely if the pressure drop at the venturi 53 indicates that the flow is falling below. a certain predetermined minimum value. This may be about 50% of the normal flow rate, or it may be set for the minimum rate of hot catalyst flow necessary to provide the heat required for vaporization of the feedstock. oil supply injected through line 13o A safety valve 55, a venturi 56 and a controller 57 play the same role in the rising branch of the U 20 just below the air injection line 36 entering the riser 22.

   In either case, the controller measures the pressure drop existing in the venturi and valve assembly, and the venturi and valve assembly can be combined into a single group, if desired.



   As an additional safety feature, flow rate controllers not shown can be provided on the oil supply line 13 and the air supply lines 36 and 37, with safety vent lines for automatically shut off the supply of water vapor or other inert gas at these points in the event that the oil supply or air supply flow rate falls below the minimum safety value to maintain fluidity, or in the event that the valves 54 and 55 close as a result of any reversal of normal operation such as a drastic decrease in the flow rate of the catalyst.

   This characteristic is particularly important at the point of oil injection, as the loss of gas velocity at this point can result in the formation of a dense phase slurry in the transfer line 30 with heavy fouling and loss. some tendency to plug the transfer line with a thick slurry of oil and catalyst.



  Appropriate check valves, not shown, may also be provided in the various air and steam injection lines to prevent any tendency for the fluidized catalyst to enter these lines in the event of a fault. line pressure,
Figure 2 shows another embodiment of this invention using the same principles of pressure differentials for catalyst density control to create catalyst flow, and U-shaped safety legs to prevent back mixing of gases. , without the use of screens to support the fluid beds of catalyst in the reactor and regenerator.



   The equipment illustrated in this drawing comprises a reactor 60 and a regenerator 61 connected by transfer pipes comprising safety branches 62 and 63 arranged at a level situated below the bottom of the reactor and of the regenerator. The reactor 60 contains a catalyst fluid bed 64, the level of which is indicated at 65. The hydrocarbon feedstock for the catalytic oracking process is injected through line 66 into transfer line 67 leading to the bottom of the reactor .



   The reactor 60, which has a cylindrical side wall and a cone-shaped bottom, has an annular scrubbing or washing space ;, disposed inside the reactor, and an extended stripper disposed below.

  <Desc / Clms Page number 18>

 of it. A dummy shell 68, having a cylindrical section at the top and a cone section at its lower part, is suspended symmetrically inside the reactor, being spaced from the side wall and the bottom so as to form a purification space. ring finger 69.

   The transfer line 67, which serves as an inlet to the reactor for the catalyst and oil, passes through the bottom of the conical section of the casing 68 with hard friction between the transfer line and the liner. The envelope is not necessarily sealed, so that a certain part of the catalyst can at this point 70 flow into the extended stripper 71 located below the reactor 60. Most of the catalyst enters the combined annular and extended stripper arrangement through orifices 72 disposed around the circumference of the cylindrical section of the casing 68. The annular washing space below the orifices 72 is preferably provided with chicanes 73, of same as the extended stripper section 71.

   A suitable source of steam or other scrub gas is supplied through line 74 at the base of scrub section 71.



   In some cases, it may be considered that either the annular scrubber device or the extended stripper, as shown above, may be sufficient on its own without the use of both devices, and it comes into play. of this invention to employ one or the other type of purification means without the other. If the annular stripper is to be used without the extended stripper, a simple vertical pipe can be provided in place of the stripper 71 to ensure the connection with the U-shaped safety branch 62, below the level of the reactor 60.



   Regenerator 61 contains a fluid bed 75 of catalyst having a level as indicated at 76, undergoing regeneration by combustion to remove carbonaceous deposits formed during the cracking process.



  The level 76 is fixed by the catalyst overflow from the fluid bed 75 into a fixed collecting cup 77. The catalyst thus withdrawn flows through a down pipe 78 into the safety branch 63, and from there into the reactor by means of of the transfer line 67 which terminates in an inverted distribution cone 79. An inverted deflector cone 80 is disposed above the entry point to cooperate with the distribution cone 79 giving a more uniform distribution. catalyst and reactive hydrocarbon through the base section of the fluid bed 64.



   The catalyst discharge pipe 78, which is shown as an outer pipe in the drawing, may, if desired, extend vertically downward through the fluid bed 75 and pass through the portion. from the bottom of the regenerator 61 instead of extending immediately through the side wall as an outer pipe.



   The washed catalyst, withdrawn from the bottom of the scrub section, passes through the safety branch 62 in the riser pipe 82 which opens directly into the bottom of the fluid bed 75, into the regenerator 61. An inverted deflector cone 83 can be disposed above the open upper end of the riser pipe 82 to give a more uniform distribution of the stream of catalyst entering at that point. The deflector cone 83, as well as the deflector cone 80, may be constructed as an inverted cone having a small central opening and having an outer diameter somewhat larger than the diameter of the catalyst stream which comes through. run into it.

   Such a construction allows part of the catalyst stream to flow continuously through the cone so that there is no dead space above it, and deflects the remainder circumferentially. so as to give a much more uniform distribution of catalyst and gas across the cross section of the vessel than is achieved by allowing the current to pass directly through the fluid bed without any distribution grids or deflection baffles.



   The primary control of the flow of catalyst in circulation is obtained in this equipment, as in the case of the variant explained previously in relation to figure 1, by adjusting the density of the catalyst suspension in the riser pipe 82 leading to the regenerator. 61. As in

  <Desc / Clms Page number 19>

 In the preceding example, this can be accomplished by supplying varying amounts of air from the main supply of air from a fan 84, to the base of the riser 82 through ducts. air injection 85, while the remainder of the regeneration air supply is introduced into the regenerator through a separate line 86.



   The various pipes and safety branches of this model function essentially in the same manner as that described above in Fig. 1. The top pressure in the reactor 60 is determined by the back pressure acting on the reactor. recovery system, not shown, and the product vapors are removed from the top of the reactor by means of a line 88 after passing through a suitable gas-solid separator 89 to remove the entrained catalyst and return it through. suitable immersion branches in the fluid bed 64. The pressure differential between the reactor 60 and the regenerator 61, if any, is maintained at a constant desired level by a pressure sensitive device connected at the top of the two containers and controlling the opening of a regulating valve 90.

   The control valve 90 is placed in a chimney 91 which removes the spent regeneration gases from the regenerator 61 after they have passed through a gas-solids separation device 92 to extract the entrained catalyst therefrom. Heat 93 may be disposed in chimney 91 to preheat the hydrocarbon feedstock introduced through conduit 94, by indirect heat exchange with the flue gases from the regenerator.



  In certain circumstances it may be desirable to employ a so-called double tubular sheet heat exchanger at this point instead of the single heat exchanger shown schematically in the drawing. Such an exchanger prevents mixing of the heat exchanger fluids during a fault at the point where the tubes pass through the collector tubular sheet. For this purpose, a narrow space open to the atmosphere is provided between two separate tubular sheets supporting the tubes, one tubular sheet serving to close the collector space, the other to close the exchanger space.



   As a safety control device for operations, a venturi flow indicator 95 may be provided in safety branch 62 receiving spent catalyst to shut off the main air fan 84 and thereby interrupt the flow of air. air in the event that the average catalyst flow rate falls below a predetermined minimum value.



  A similar venturi indicator 96 may be disposed in the safety branch 63 for the regenerated catalyst, to control a valve 97 provided in the oil supply line 94, and to shut off the oil supply to the line. injection 66 in the event that the catalyst flow in the safety branch 63 falls below the predetermined minimum.



  The air flow controller 98, which determines the relative amount of the regeneration air supply introduced through lines 85 and 86 to the base and top of the riser 82 respectively, can be operated by a controller. temperature placed in the regenerator 61, as shown here, or in the reactor 60 as is the case in the variant described in Figure 1.



   Note that different elements of these two models can be used in combination. Thus, for example, either the annular stripper or the external regenerated catalyst pipe of figure 2 can be used instead of the corresponding elements of the model shown in figure 1, and different advantages can be found in operating without distribution and support grids 32 and 42, in a manner similar to that described in Figure 2.



   As a specific example of densities and pressure differentials suitable for operation in such a system, a catalytic cracking plant designed according to Figure 1 and using a synthetic silica-alumina catalyst in the form of spray-dried microspheres and having a distribution of dimensions as follows. g 7% at 0-20 microns, 15% at 20-40 microns, 66% at 40-80 microns and 12% at 80 microns and more, may present a pressure at the top of the reactor of 9.8 pounds per inch square, a bed density in the reactor of

  <Desc / Clms Page number 20>

 28 pounds per cubic foot for a normal bed depth of 25 feet,

   a catalyst density of 32 pounds per cubic foot for a depth of 16 feet in the stripper section 17, a pressure of about 18.1 pounds per square inch at the inlet 19 for the scrub gas, and a density of 38 pounds per cubic foot in the safety branch 20 as well as in the safety branch 27 and the down pipe 24;

   under these conditions, the pressure on top of the regenerator can be about 11.2 pounds per square inch, with a catalyst density of 23 pounds per cubic foot for a depth of 13 feet in the fluid bed 15, giving a 13.3 pounds per square inch pressure at the grid 42 and 16.3 pounds per square inch pressure at the air injection line 36 when the density of the suspension in the riser 22 is of about 17 pounds per cubic foot and that the density of the diluted slurry in transfer line 30 is about 6 pounds per cubic foot.

   Under these conditions the pressure drop across the fluid bed 11 in the reactor is about 4.8 pounds per square inch, the pressure drop across the fluid bed 15 in the regenerator is about 2. 0 pounds per square inch and the catalyst level 50 in discharge well 25 is approximately 7 feet below level 16 of bed 15. The height of the safety legs below the injection line d The oil 13 and air injection line 36, above the bottom of the U-shaped sections 27 and 20 respectively, is about 15 feet, which is enough to create a fluidistatic safety pressure of about 4 pounds per square inch in each of these.

   A safety limb height as small as about 7.5 feet to give a safety effect of a 2 pound pressure could be a sufficient safety factor in systems of this type where pressure fluctuations are very high. weak. In other cases, correspondingly longer provided safety legs may be employed to allow greater pressure fluctuations, especially in operations performed at higher vessel pressures or with a greater difference between the pressures. operating pressures in the two vessels.



   Thanks to the U-shaped safety branches described, these designs are both suitable for use in catalytic cracking plants operating at surface gas velocities considerably higher than those employed in current commercial practice. The specific data shown in the example above is suitable for operation with the preferred speed range of about 2.5 to 3.0 feet per second in both the reactor and the regenerator.

   This process will also work equally well at low speeds on the order of the minimum catalyst aeration speed, which may be between about 0.1 and 0.5 feet per second, and higher speeds may be. employed, up to about 6 to 8 feet per second, which is considerably greater than the rate of free fall of the catalyst particles employed.



   Although this process has been described using a specific silica-alumina catalyst, other cracking catalysts may be suitable as well. In general, the figures given here relate to an operation carried out with ordinary cracking catalysts. With catalysts or solids having a much higher or much lower density, the given densities and speeds will vary but the same principles of fluidistatic pressure control will still apply. The temperatures, the average flow rates and the severity of the cracking employed form no part of this invention and are not limited to it.



   Although the above disclosure has described, by way of example, the specific application of this invention to systems for the catalytic cracking of hydrocarbons, it will be appreciated that the same basic principles can be employed with advantage throughout. series of other applications. Thus, the invention is suitable for the hydroformation of naphtha, and other important uses can be found in the field of the catalytic conversion of gases or vaporized materials wherever it is necessary to regenerate a finely divided solid catalyst at high levels. frequent intervals, by oxidation, reduction, sulfurization or other special treatment, after

  <Desc / Clms Page number 21>

 that it is exhausted during the conversion period.

   One example of the many non-catalytic processes in which such a system can be employed is the thermal cracking of heavy carbonaceous materials such as reduced crude product, using an inert solid recycle stream as a heat carrier. Another example exists in the treatment of a stream of gaseous or liquid hydrocarbons mixed with silicate gel, activated charcoal or other solid adsorbents suitable for the removal and recovery of a desired component from a compound. such mixture by selective adsorption.

   Another example is the distillation of oil from oil shale, where the spent shale can be roasted in a stream of oxidizing gas, in a separate vessel, and returned through the cycle to the distillation vessel to provide heat. for the distillation of a new quantity of oil from the fresh shale introduced Other uses may include the carbonization of coal, the roasting of mines, the reduction of metal oxides or the transmission of heat between streams of fluids by means of solids.



   Referring now in detail to Figures 3, 3a, 4 and 4a of the drawings, where like parts are indicated by the same reference notations, the system illustrated therein essentially comprises a retort 110 for the product. crude, a hydroformer or a naphtha reformer schematically represented by an element 130, a product fractionator 140, a cracking phase schematically illustrated by an element 150, and an apparatus 190 for the possible filtering of fuel oil.

   The role and the coaction of these elements will be explained immediately by taking as an example the refining of a crude oil of average specific weight of the Arabian Quatar type in a refinery with a capacity of about 100,000 barrels produced crude per day. It should be understood, however, that the system can be employed for the refining of different types of crude oils on a larger or smaller scale in a generally analogous manner.



   In operation, crude oil is pumped from a line 101 by means of a pump 103, through line 105 and through heat exchangers 106, into a heating coil disposed in a furnace 107 where it is heated to a suitable temperature. to vaporize a significant portion of the oil. This so heated is sent through line 109 to a lower portion of retort 110 into which it can enter at a temperature of about 750 to 800 F and a pressure of about 40 to 70 pounds per. square inch.

   The retort 110 may be provided with a series of horizontal cap plates 112 for the bubbles to improve the fractionation of the feed in a conventional manner. The reflux may be carried out using partial condensers 114 arranged. internally or externally at the top of the retort 110. For use in the present example, the retort 110 can be operated so that three distillate streams and three still bottoms are produced as follows:
All crude constituents boiling below about 250 F are removed together as a stream of light virgin naphtha vapor, through the top in line 116, at a temperature of about 300- 350 F. This stream can be removed. be about 20 to 25% of the crude product loaded.

   A liquid stream of heavy naphtha having a boiling range of about 250 to 400 or 500 F, is removed through line 118 from an upper portion of retort 110 to a point below condensers 114. About 20 to 25% of the crude product loaded is recovered through line 118. Kerosene or diesel oil boiling in the range of about 400 to 700 F and amounting to about 17 to 23% of the crude product is removed through a line 1200 Le remainder of the charge, amounting to about 40-50% and consisting predominantly of constituents boiling above 700-800 F, is withdrawn as reduced crude through line 122 from the bottom of retort 110.

   The kerosene withdrawn through line 120 is normally suitable for clean kerosene or diesel oil applications without further processing, and can be sent directly to storage. The other fractions can be processed according to the present invention as will be described immediately below.

  <Desc / Clms Page number 22>

 



   The light virgin naphtha vapors in line 116 can be fed directly to a lower portion of a product fractionator 140. If desired, this vapor stream can be preheated to about 800 to 1,000 F to accommodate. to the thermal requirements of fractionator 140. This can be achieved by bypassing at least part of the vapors from line 116 through a heating coil 117 arranged either separately or, as shown in Figure 3, in a convection section 132 a reforming oven 130 which operates as described below.



   The heavy naphtha stream in line 118 can be pumped by pump 126, through line 128 at a pressure of about 900-10100 pounds per square inch, in a catalytic hydroformation phase. lytic or thermal reformation. This reforming phase can be of any conventional design well known in the art, and can include a conventional tube furnace 130 designed to allow an oil residence time of about 18 to 25 volumes of liquid per volume of reaction space. per hour (v / v / hr) at about 1,000-1.00 F and at a pressure of about 10,000 pounds per square inch.

   Under these conditions, the octane value of the naphtha can be increased from an octane number of about 20 to 30 to an number of about 70 to 80 without excessive cracking into normally gaseous hydrocarbonsa. in another way, this reformation phase can comprise a hydroformation unit as described below with reference to FIGS. 4 and 4a. The total effluent from reformer 130 is sent to substantially the pressure and reformation temperature through line 134, provided with a pressure relief or safety device such as a valve or a venturi 136. , in a lower part of fractionator 140 at a point located near the feed point of pipe 116.

   As a result of the change from pressure to fractionator pressure of about 5 to 15 pounds per square inch, the naphtha is almost completely vaporized as it enters the lower portion of fractionator 1400
The reduced crude product from line 122 can be sent directly to the bottom of fractionator 140, at substantially the temperature it is at when it is removed from retort 110. Line 122 enters the chamber. fractionator 140 at a point above the inlet or supply points of lines 116 and 134. In this way, the vapors supplied by lines 116 and 134 pass upwards through fractionator 140 at the top. against the reduced crude product flowing downward to free the latter of vaporizable constituents.



  This effect and the operation of fractionator 140 will be described in more detail later.



   At this point, it can be seen that a side stream of hydrocarbons from the gas oil range amounting to approximately 45 to 60% calculated on the crude product and boiling between approximately 600 and 1,000 to 1,100 F, which is suitable as a feedstock. catalytic cracking, can be removed from an intermediate section of fractionator 140 by a gas oil reflux system, comprising a pump 143 and lines 142 and 144, and sent via line 146 to a catalytic cracking stage 150. This cracking phase is preferably designed as described in Figures 1 and 2, but any other conventional cracking system, intended to convert hydrocarbons from the gas oil range into oils with a lower boiling point, in particular from the fuel range of motor, can be used.

   Continuous or batch operation can be employed in fixed bed, moving bed, "fluid" or suspension systems. The heat required for cracking can be supplied as pre-heat of the process substances and / or as the sensible heat of the exothermically regenerated catalyst, or in any other customary manner.



  Catalysts of the frozen type or synthetic or natural modified clay, such as activated montmorillonite clays, silica-alumina compounds, silica-magnesia and other conventional cracking catalysts can be employed at temperatures of about 800 to 1,000 F and pressures ranging from about atmospheric pressure to 25 pounds per square inch, all in a manner known per se.

  <Desc / Clms Page number 23>

 



   The total hydrocarbon effluent from the cracking phase 150 is sent to substantially the cracking temperature of, say 800 to 900 F approximately through a line 152 to the lower part of the fractionator 140, preferably to a point intermediate between the feed points for reduced crude product on one side, and virgin and reformed naphtha on the other side. If the cracking phase 150 is operated at a high pressure, the pressure may, under the action of a valve 152a, change to the pressure of the fractionator.

   In most cases, about 95 to 100% of the cracked or cracked material enters fractionator 140 in the vapor state to enhance the washing action of the vapors supplied by lines 116 and 134, while any non-vaporized component of the cracked material is in turn subjected to washing by the vapors which are introduced through lines 116 and 1340.
As shown in Figures 3a and 4a, fractionator 140 comprises a lower wash section A and an upper combined fractionation-absorption section B. Both sections are provided with suitable means to improve the countercurrent contact between the descending liquid and ascending vapors.

   For cleaning purposes a disc baffle arrangement has been found to be most effective and is shown schematically for section A by elements 141. Section B is shown as containing a number of cap plates 154 for the bubbles to enhance the efficiency of the combined fractionation and absorption process.



  Sections A and B can operate as follows.



   The washing section A receives, in addition to the vapor and liquid streams supplied by the lines 116, 134, 152 and 122, a liquid supply, at the top, comprising a part of gas oil removed from the bottom of the section. B via line 142 and supplied to section A via line 144. This gas oil is fed to section A to provide control of reflux and heat removal in that section to achieve the desired final purification. diesel. All the heat required for washing and fractionation in fractionator 140 is preferably supplied as sensible heat from the hydrocarbon streams entering section A to maintain a temperature of, say, 820 - 830 F or so in the tank. lowest part of section A.

   The vapors rising through section A frees the gas oil, reduced crude product, and downwardly flowing cracked liquid products of substantially all of their distillable components, and this vapor mixture then passes to a temperature of about 700 to 750 F in fractionation-absorption section B to be processed as described later.



   The reduced crude product from line 122, which may contain up to about 75% gas oil suitable for feeding the catalytic unit, is backwashed and heated by the cracked vapors to, say. , about 875 F and at a pressure of 8 pounds per square inch, and followed by the virgin naphtha vapors at, say, 800 F or so, and by the reforming product at about 1.025 F. The effect of The partial pressure of the other streams and their degree of heat are sufficient to cause vaporization of the gas oil constituents of the reduced crude product and to provide the required heat.

   The net effect of the process in Section A is then (1) a bottom stream of unmelted fuel oil amounting to about 10 to 15% of the crude product and containing about 85 to 90% of reduced crude product burnt, about 8 to 9% tar from the reformer, and about 0.3% heavy oil in the slurry state from the catalytic operation, all mixed automatically so that said stream can be melted with about 50% d mixed light diesel oil for correcting the viscosity of the fuel oil, and (2) vapors containing all of the distillate products to be obtained from fractionator 140, and leaving section A from above at about 800 F <,
A heavy material containing all the non-distillable constituents of the crude product charged, and fractions converted in phases 130 and
150,

   is collected at about 820-830 F in the bottom area of Section A from where it can be removed via line 147. If desired, the temperature

  <Desc / Clms Page number 24>

 re at the bottom of section A can be reduced to, say, about 700 F by re-cycling the heavy tails coming from line 147, by means of a pump 149 through a refrigerator 153 and a line 155. The Cooling of the bottoms may be desirable to prevent cracking and coking of heavy liquid products.

   The combined reduced crude product, amounting to about 12-18% based on the crude product, can be collected through line 181 for further processing as will be shown below.
Under the conditions of the present example, about 10,000 to 1,500 molecule-grams / hour of hydrocarbon vapor will be available for washing in section A about 100 to 200 molecule-grams / hour of liquid. This favorable vapor-liquid ratio results in producing a strong cleaning effect in section A.

   The number of baffles and the dimensions of section A depend largely on the character of the raw product loaded and the products desired. For the purpose of this example, this section may be approximately 12 'in diameter and 30' in height, and is provided with of 7 sets of disc type contact devices.



   Vapors passing upward through section B of fractionator 140 are refluxed gas oil through line 144 and refrigerator 1450 Gas oil can be withdrawn through line 142 at about 650-750 F and can be split. in three streams, namely (1) reflux through line 144 as described, (2) a liquid product amounting to about 2.5 to 3% of the crude product for blending with heavy diesel oil is dissolved. - directing through a pipe 148 and a refrigerator 161 to the storage, and (3) approximately 50 to 55% with respect to the crude product of the catalytic feedstock,

   which are sent directly to the cracking section by a pump 143 through a line 1460 This gas oil contains all of the virgin fraction at 600-1.050 F and all of the gas oil in the cycle, and the gas oil is sent from there to extinction.



   Turning now to section B of fractionator 140, it will be noted that in most conventional catalytic cracking operations and similar refining operations, products are fractionated at low pressures to produce gas and gas. Low Pressure Distillate The gas stream contains appreciable amounts of gasoline constituents and it is therefore necessary to compress, absorb and refract this stream to recover its gasoline constituents. This can be avoided by combining both fractionation and low pressure absorption in the upper section A of fractionator 14.

   For this purpose, one of the pumping agents normally used simply for the removal of heat and returned to a point near their extraction, can be employed as an absorption agent by returning it to a point substantially above. from the one from which it is removed.



   The operation of the upper part of section B may be briefly described here as it contributes to the achievement of the essential advantage claimed for this invention which is to make small refineries fully competitive with respect to this invention. towards the big ones.



   The vapors leaving the gas oil fractionation section and entering the heating oil extraction part contain heating oil, total naphtha and gas. In the section immediately above the exhaust line 156 for the heating oil, these vapors, which can be about 510-520 F, are cooled by contact with cold heating oil. , at about 130 F, entering the tower through line 1680 The heating oil is condensed off and falls with the cooled portion entering through line 168, and both are withdrawn from the tower through line 156. The cold heating oil in line 168 is saturated with the like C5, C4 and C3 to form a fatty oil.

   The light fractions are carried away by the ascending naphtha and gas values, thus increasing the concentration of said fractions in the condensation zone of the naphtha and increasing their absorption in the naphtha.

  <Desc / Clms Page number 25>

 



   Heater oil passing through line 156 is withdrawn at about 480-500 F, and most of it is cooled in a refrigerator 158 and returned to the top of section B through pump 160 and line 162 as lean absorber oil, the remainder being purified in stripper 157 and recovered through line 178 in the form of flux and product, the vapors being returned through line 159.



   The vapors entering the part of section B which is located above the inlet of line 168, therefore consist of naphtha fractions and gas fractions as normally encountered plus a quantity abnormal of C4, C5 and C6 which is absorbed as previously explained at the top of the section Bo These vapors and the heating oil are fractionated by a back-reflux pumped through line 176, and the fractionated vapors can be taken from the tower at a pressure of about 5 pounds per square inch and a temperature of about 215 F through line 166, and can be compressed by a single stage blower 169.

   The naphtha with an excess of light fractions is condensed in a condenser, 172, freed of the excess light fractions in a stripper 173 and, except for the part returned as reflux through line 176, sent through line 170 to storage as final product
The gas and the light fractions are returned to the absorption part of section B by a line 177 where the countercurrent absorption of the desired light fraction by the lean oil takes place, as previously described. . The number of plates between lines 156 and 162 is preferably increased by about 10 to 15 over that of normal fractionator models. Using this technique, it is possible to absorb essentially all of the C5 "fractions of the gas entering the top section of fractionator 140.

   Further, up to 75% of the C4 constituents can be absorbed. The gas leaving the top of fractionator 140 through line 164 is thus stripped of its valuable gasoline constituents and can be sent directly for use as fuel. A naphtha or gasoline uptake can be recovered via line 1660. Such gasoline can be withdrawn at temperatures of about 120-150 F despite the low pressure employed.

   If the raw product is such or the distribution of the products so requested, it is possible to operate section A at pressures below those of section B, say 5 to 15 pounds per square inch, to allow a reduction of the crude product to much lower bottom products, while at the same time obtaining high recoveries of C4 in section B while maintaining a higher pressure.



   The gasoline fraction is not contaminated by the pumping agent, as described above, and is condensed either in the tower by back-pumping naphtha or in an external condenser, part of the naphtha being returned to the tower. the tower for splitting. The number of plates to be provided in section B will depend on the type of raw product loaded and the products desired.

   For the purposes of the present example, two plates may be employed between the inlet of reduced crude product and the extraction of gas oil, four between gas oil and heating oil, five for the removal of fractions. gas heating oil, three for naphtha-heating oil fractionations, and ten to fifteen plates for absorption of light constituents in the top part of the Bo section
As shown in the drawings, the final products can be recovered from fractionator 140 as follows. Peak 400 F gasoline, amounting to about 50 to 60% of the crude product and having an octane number of about 75 to 85, can be sent to storage via line 170.

   About 20 to 30% gasoline of the crude product can be recirculated by pump 174 via line 176 to section B to serve as a reflux.



  Light diesel oil or final heating oil can be recovered via line 178 in an amount of about 0.5 to 1.0% of the crude product A mass of heavy diesel oil, representing about 2 to 3% of the product raw, can be obtained via pipe 1480

  <Desc / Clms Page number 26>

 
Returning now to the combined reduced crude type bottoms, withdrawn through line 181, they can, if desired, be mixed with gas oil or lighter fractions supplied through line 183, to adjust. their viscosity in order to meet the requirements ,, The tails can then be cooled to about 200 to 500 F in a refrigerator 185, and sent through line 187 to a filtering arrangement 1900 Conventional sand filters,

   rotary filters or porous bonded ceramic filters can be employed to remove from the combined residue any suspended or slurrying solid particles, such as coke, catalyst carried off from cracking phase 150, and the like. The solids removed in the filter arrangement 190 can be discharged via line 192 or sent to a catalyst recovery device (not shown). A fuel oil grade residue is recovered via line 194.



  The combined filtering of thermally cracked tar, catalytic sludge and crude product residue, of this kind, avoids the "compatibility" problems which arise in the usual mixing of such materials, and allows the recovery of a material. relatively low deposit fuel oil,
The system illustrated in Figure 3 allows various modifications.



  As indicated above, one or more phases of coking or breaking the viscosity of the reduced crude product can be intercalated. Some of the effects envisaged in the case of such coking phases, can be achieved by operating the bottom part of the section. A fractionator 140 under viscosity breaking conditions, for example at 875 up to 950 F. Instead of, or in addition to, this type of operation, the reduced raw product from line 122 can be subjected to a conventional treatment of viscosity breaking or catalytic or thermal coking.



  The total effluent from such a coking phase, with or without the coke formed, can be sent through line 122 to an upper part of section A of fractionator 140 for treatment there as described above. Similarly, some or all of the tails of line 147 may be subjected, for example, in unit 182 to viscosity breaking and / or coking to produce heavy residue and coke to be sent from a stripper 184 via lines 186 and 181 to filter device 190, and lighter products which can be returned from stripper 184 via line 188 to a middle or middle part of section A for washing and to fractionation as described above.



   Catalytic, rather than thermal, reforming can be employed in reforming step 130 using conventional catalysts such as oxides and sulfides of Group V, VI or VIII metals, preferably supported by a suitable vehicle, temperatures of about 850-1100 F and pressures ranging from atmospheric pressure to about 400 pounds per square inch in the presence or absence of external hydrogen, all in a manner known per se. Other refining treatments such as bauxite treatment, clay treatment, etc. can follow the reformation phase 130, provided that most of the hydrocarbon effluent from these phases is supplied to the fractionator. 140 as described above.



   Although various catalytic cracking systems can be employed as the cracking phase 150, a "flowable" catalytic cracking involving the continuous production of cracked effluent and a continuous flow of catalyst between the cracking and regeneration phases, as described with reference to Figures 1 and 2, is particularly suitable.



   Referring to Figures 4 and 4a, the illustrated embodiment is very similar to that described above and illustrated in Figures 3 and 3a. The operation of the system is the same as that described in connection with Figures 3 and 3a, with this variant that the thermal or catalytic reforming phase 130 is replaced by a catalytic hydroformation unit,
If we study the operation of this unit, we see that the heavy naphtha stream can be pumped by a pump 126 through a pipe.

  <Desc / Clms Page number 27>

 128, at a pressure of about 600 to 900 pounds per square inch, to a catalytic hydroformation phase 130.

   This hydroformation phase can be of any conventional model well known in the art, and is preferably of a model suitable for a platinum training operation (platforming) or a regenerative hydroformation operation of the type specified above, For the present example, a conventional catalytic reforming operation carried out at about 800 to 1000 F, a pressure of 50 to 1000 pounds per square inch, a nominal oil residence time of about 0 is contemplated. , 5 to 2.5 hours, calculated as cold oil in contact with the catalyst, and hydrogen recycle flow rates of about 500 to 70,000 standard cubic feet per barrel of oil, using a catalyst, such as than molybdenum supported by alumina, in a fixed bed operation.

   In addition to the virgin naphtha supplied to the hydroformation phase 130 through line 128, a hydrocarbon stream of a similar boiling range, derived from the fractionator section of tower 140, can be supplied to the hydroforming phase 130 via line 129 in any desired proportion. Under the conditions specified, the octane number of the naphtha can be increased from about 25 to 40 to about 80 to 100, without excessive cracking into normally gaseous hydrocarbons and carbon.

   A series of hydroformation reactors can be provided alternating between current and regeneration cycles, oxidizing gases, such as air, water vapor, flue gas or mixtures thereof, being used in the regeneration cycle to remove carbon from the catalyst, all in a manner known per se, Of course, any type of continuous operation such as fluid, moving bed or suspensoid operation can be used in wherein the catalyst is continuously circulated between reaction and regeneration vessels.



  If desired, external hydrogen can be supplied to the hydro- formation phase 130 via line 1310.
The total effluent from hydroformer 130 is sent through line 134, via heat exchanger 117, to stabilization tower 200, in which hydrogen and normally gaseous hydrocarbons such as C1 - C3 hydrocarbons are separated from the normally liquid hydroformate at temperatures of about 80 to 120 F and pressures of about 100 to 10,000 pounds per square inch.



   Gases containing about 40 to 90% hydrogen and amounting to about 1,000 to 9,000 standard cubic feet per feed barrel are withdrawn from stabilizer 200 through line 202 and passed through a recycle booster or booster 204 and a heater 206. An amount corresponding to about 500 to 70,000 standard cubic feet of hydrogen per feed barrel may be supplied via line 208 to hydroformer 130 at substantially the temperature and pressure thereof.

   The remainder of the gases in line 202 or any desired portion thereof may be passed through line 210, via a booster 211 and lines 212 and / or 214, to hydrogenation phases 216 and / or. or 218 which can be used to improve the cracking characteristics of the cracking recycle feed supplied from tower 140 through line 142, or to improve the quality of light gas oil recovered from tower 140 via line 178, as will appear more clearly below. Any excess gas can be removed through line 203 and sent through line 203a to the bottom portion of stripper-fractionator 140 as additional scrubbing agent, or used as flue gas in the system.

   By operating in this manner, for example, the density of the recycle gas oil can be increased from 2 to 15, and the feed can be made as desirable as the virgin feedstock for the catalytic cracking.



   The hydroformate separated in stabilizer 200 is withdrawn through line 220. This substance can be mixed without further treatment in tank 222 with finished or completed gasoline recovered from tower 140, as will become clearer below. The hydroformate thus recovered can amount to about 10 to 25% of the crude product.

  <Desc / Clms Page number 28>

 



   As shown in Figure 4a, part of the side stream of gas oil hydrocarbons withdrawn from the fractionator via pump 143 and line 142, can be sent through a hydrogenation phase 2160.
For example, all the heavy gas oil which is not to be returned as reflux to fractionator 140 through line 144, can be supplied by pump 143 to the hydrogenation stage 216 and hydrogenated there with tungsten sulphide. nickel as a catalyst at a temperature of about 750 to 800 F, a pressure of about 20,500 to 3n000 pounds per square inch,

   a liquid flow of about 1 to 2 V / V / Hr and a hydrogen feed rate of about 2,500 to 3,000 standard cubic feet per barrel of oil in a fixed bed operation. Under these conditions, approximately 1,000 to 10,200 standard cubic feet of hydrogen per barrel is consumed in the hydrogenation process.



  The API density of the heavy gas oil in the cycle is increased from about 10 to 12 units and substantially all of the sulfur is removed. The hydrogenated gas oil can then be recirculated through line 146 as described above.



   In the embodiment shown in Figures 4 and 4a of the drawings, the final products from tower 140 can be recovered in a slightly different manner from that shown in Figures 3 and 3a above, as follows. Gasoline having a peak temperature of 400 F, amounting to about 35-45% of the crude product and having an octane number of about 80-90, can be sent via, line 170 to a phase of. conventional finishing 221 and then in a tank 222 to be mixed there with the hydro-format supplied via line 210. The gasoline can be recirculated by a pump 174 via line 176, to a section B, for use. reflux A light or final heating diesel oil can be recovered via line 178 in an amount of about 0.5 to 1% of the crude product.



  A heavy diesel oil charge, representing about 2-3% of the crude product, can be obtained via line 148. A fraction of naphtha somewhat heavier than that of gasoline in line 166, and having a range of boiling of, say, about 200 to 400 F, can be removed from an intermediate point between the discharge point of line 156 and the feed point of line 168, to be sent via line 129 to hydroformer 130 as described above. The quantity of naphtha thus removed can vary between 0 and 45% approximately of the crude product. depending on the proportion of virgin naphtha supplied directly to hydroformer 130 via line 128.



   Another embodiment of this variant of the invention is illustrated in Figure 5 which, for the sake of simplicity, shows only those elements which differ from those of Figures 4 and 4a from the point of view of design and / or operation, and in addition, some of the lines of Figures 4 and 4a to clearly show how the equipment illustrated in Figure 5 is to be associated with the combination group described with reference to Figures 4 and 4a. In essence, the elements shown in Figure 5 are intended to replace hydroformer 130, stabilizer 200, and hydrogenators 216 and 218 of Figures 4 and 4a.



   Referring now in detail to Figure 5, the system illustrated therein essentially comprises three catalytic hydroformers 300, 302 and 304 connected in series, a hydroclarification phase 320 and a stabilizer-fractionator 330. The role of these elements in a combined group of the type illustrated in Figures 3, 3a, 4 and 4a, will be described immediately in connection with a platinum training operation (platforming) performed in hydroformers 300, 302 and 304. It should be understood. , however, that other types of catalytic hydroformation can be carried out in this system in a substantially analogous manner.



   In operation, naphtha having a boiling range of about 150 to 400 F can be sent through line 128 into hydroformer 300 via heater 307, at a temperature of about 900 to 1,000 F, a pressure of about 50 to 10,000 pounds per square inch and in a proportion of, say, about 15 to 45% of the crude product. This naphtha can be supplied from retort 110 via line 118 and / or from tower 140 via the driving

  <Desc / Clms Page number 29>

 129 in any desired proportion, as described with reference to Figures 3, 3a, 4 and 4a.

   From about 500 to 70,000 standard cubic feet per barrel of recycle hydrogen is supplied to hydroformer 300 through line 306 at hydroforming pressure and temperature of about 900 to 1,200 F. The hydroformers 300, 302 and 304 contain a platinum group catalyst such as platinum or palladium supported on alumina pellets preferably arranged in fixed beds. The naphtha-hydrogen feed passes in series through hydroformers 300, 302 and 304 and is reheated between phases in heaters 308 and 310 to maintain similar temperatures in all three phases.

   Suitable platinum forming conditions, to be maintained in hydroformers 300, 302, and 304, include temperatures of about 800 to 1,000 F, total pressures of about 400 to 10,000 pounds per square inch, and recycle rates of. hydrogen of about 3,000 to 12,000 standard cubic feet per barrel. Under these conditions, the octane number of naphtha is improved anywhere from 30 to 60 points and the sulfur is removed almost completely. The catalyst retains its activity for many months, after which it can be replaced with fresh or regenerated catalyst, without the need for frequent continuous or periodic regeneration.



   The total effluent from the hydroformation or platinum formation phases passes through line 312 to a hydroclarification phase 3200 Simultaneously, the hydroclarification phase 320 receives via line 119 all or any desired part of the gas fraction light oil recovered from tower 140 before hydrogenation in hydrogenator 216. The combined feed is treated in hydroclarification phase 320 with a hydroclarification catalyst such as oxides or sulphides of groups V metals, VI and VIII, at temperatures of about 500 to 800 F, pressures of about 50 to 10,000 pounds per square inch, and passages of about 0.5 to 5 V / V / Hr.



  Under these conditions, substantially complete conversion of the undesirable constituents of the light gas oil fraction, such as sulfur compounds, etc., takes place by direct hydrogenation or by transfer of hydrogen from the formation product to platinum, without affecting appreciably the quality of the platinum forming product. As a result of the high hydrogen partial pressure, the deactivation of the catalyst by carbon deposition is relatively slow, eliminating the need for frequent continuous or periodic regeneration. Fixed-bed operation is therefore suitable. for the hydroclarification phase 320. However, any conventional continuous or periodic regeneration system, comprising a fluid or moving bed operation, can be employed, if desired.



   The effluent from the hydroclarification phase 320 is sent through line 322 to an intermediate section of a conventional stabilizer-fractionator 330. This tower operates so that a heating oil or diesel bottoms fraction of high quality can be recovered and sent to storage via line 332, the platen forming product having a final boiling point of about 400 F can be sent via line 334 to the mixing tank 122, and gases containing about 80 to 95% hydrogen and amounting to about 5,000 to 12,000 standard cubic feet per barrel are recovered from above through line 336. About 500 to 10,500 cubic feet per barrel of gas from the line 336 can be disposed of as excess.

   The remainder is returned to the cycle using a stimulator 340, via lines 338 and 342 and a heater 344, to line 306 and hydroformer 3000
Rather than simply sending the light gas oil fraction from line 119 to the hydroclarification phase 320, any desired portion of the recycle feed from line 146 can be used for this purpose. In this case, the fraction of oil recovered through line 332 can be returned in whole or in part to line 146 to serve as an improved recycling charge for cracking phase 150.



   In all other respects, the operation of the entire combined group may be substantially as described with reference to Figures 1 and 1a, including the operation of retort 110, of tower 140,

  <Desc / Clms Page number 30>

 of the cracking phase 150, of the coking phase 182, etc.



   Still another embodiment of the variant of the invention is illustrated in figure 6 which shows a schematic flow plane of a complete combined unit similar in various aspects to that shown in figures 3, 3a, 4 and 4a, the corresponding elements, although illustrated in a simplified manner, being indicated by the same reference notations.



  An important additional element is a high pressure stripper 435, the role of which will be described immediately.



   Referring now in detail to Figure 6, crude oil is charged via line 109 into retort 110 and processed there substantially as described with reference to Figures 3 and 4. A kerosene or gas fraction Ready-to-use light oil can be recovered via line 120. Light naphtha is withdrawn via line 116, heavy naphtha is withdrawn via line 118 and reduced crude product via line 1220 In addition, a fraction of gas oil can be removed. be recovered by a pipe 420 and sent in whole or in part to a cracking phase 150 to serve as a virgin cracking charge. Virgin gas oil can be recovered as a finished product via line 421.

   Any desired proportion of light and heavy naphtha can be recovered via lines 416 and 418 as a finished product or sent directly to the base of tower 140 as described with reference to Figures 3 to 50.
According to a preferred variant of this embodiment of the invention, the heavy naphtha coming from the retort 110, representing about 15 to 20% of the crude product, is sent to the pipe 128 and the hydroforming phase 130. The latter can work substantially as described for phase 130 and hydroformers 300, 302 and 304 of FIGS. 3 and 5, with recycle hydrogen being added via line 406, as will appear more clearly below.



   The total effluent from hydroforming phase 130 is sent to substantially the temperature and pressure of phase 130, or about 800 to 1,000 F and about 700 to 800 pounds per square inch, through line 432. to a bottom part of a high pressure stripper 435.

   This stripper, which can be of any conventional model, having a lower section C with discs and rings and an upper section D with trays for the bubbles, is maintained substantially at the pressure of the reformation phase 1300
Reduced crude product from line 122 is sent by pump 437 through line 439, at the pressure of stripper 435, to the top of section C of stripper 435 which is designed so that the temperature varies by approximately 600 to 800 F at the bottom to about 500 to 750 F at the top of the stripper 435. Under these conditions, the reduced crude product supplied by the line 439 is freed of its gas oil and its constituents at low boiling point in a proportion of approximately 5 to 35% of the reduced crude product,

   depending on the type of raw product considered. The reduced, washed crude product, then amounting to about 25 to 45% of the original crude product, is withdrawn through line 441, fitted with a safety valve 443, and sent to an upper part of the tank. section A of tower 140, to be further treated there by cracking vapors coming through line 152 from cracking phase 150, in substance as described with reference to Figures 3 to 4a.



   A combined overhead product containing the reduced crude product wash, consisting primarily of gas oil, total hydroformate and total hydrogen, is withdrawn from stripper 435 via line 445, cooled in refrigerator 447 to at about 80 to 120 F and passed through a gas / liquid separator 449 with virtually no reduction in pressure other than that caused by cooling.

   The gas separated in separator 449, amounting to about 500 to 50,000 standard cubic feet per barrel and containing about 40 to 95% hydrogen, is removed through line 406, reheated in heater 451 to about 900 to 1,200 F and returned to the hydroformation phase 1300 Excess gas can be released through line 407 to be used either as flue gas or, in tower 140, as an agent

  <Desc / Clms Page number 31>

 additional washing.



   The total liquid separated in separator 449 can be sent through lines 453 and 455 fitted with a safety valve 457, to fractionation section B of tower 140 for fractionation there substantially as described with reference to the figure. 4a. If desired, some or all of the liquid in line 453 can be sent to a separate fractionator 460 after releasing pressure in valve 4590 High quality hydroforming material can be collected from above from tower 460, and be recovered via line 462 to be mixed in tank 222 of Figure 4a with finished gasoline from tower 140.



  A fraction of gas oil can be withdrawn from tower 460 as a tail product via line 464 to be combined with gas oil from line 142 and then processed as described with reference to Figures 3 and 3a.
In a similar fashion to what has been explained with reference to Figures 4, 4a and 5, the virgin naphtha from lines 116 and / or 118 may be supplied in whole or in part directly to the base of tower 140 to serve as The reduced crude product wash, and corresponding amounts of similar fractions from tower 140 may be employed as feed for hydroformer 130, as will immediately be apparent to those skilled in the art. Also, any desired portion of virgin light naphtha from line 116 can be sent to reforming stage 130.



    CLAIMS.



   1. A method of controlling the flow of a finely divided solid through a fluidized bed maintained in a contact vessel, which comprises introducing the solid from bottom to top into said bed through the ascending branch of a. open U-shaped conduit, opening into said receptacle and extending below it, and the extraction of the solid from said bed from top to bottom by the descending branch of a second conduit in U-shape opening into said container and extending below it,

   supplying a fluidizing gas to the lower part of each of said U-tubes in an amount sufficient only to maintain fluidification of the solid in said bed and with a maximum density of the mass of fluidized solid which s' found there by adding a fluidizing gas in the descending branches of each of said U-shaped conduits sufficient to maintain the density of the column of fluidized solid in said descending branches and by adding a fluidizing gas in the ascending branches - Your of each of said U-shaped ducts sufficient to maintain the density of the fluidized solid column which is therein less than the maximum density prevailing at the bottom of the U-shaped duct and also less than the density of the column located in the descending branch of the same U-shaped duct,

   thereby establishing a pressure differential between the descending branch of said first U-shaped duct and the ascending branch of said second U-duct, which causes the fluidized solid to flow upwardly through the duct. ascending branch of one of the U-shaped ducts, in and through the bed, and downwards by the descending branch of the other U-duct, and whereby the average flow rate is changed by varying the differen - tial of density between said two branches of the U-shaped conduits by varying the quantity of fluidizing agent supplied to the ascending branch of one of the two U-shaped conduits.


    

Claims (1)

2. The method of claim 1, wherein the ascending branch of the second U-shaped conduit of said contact container heads upward into a second contact container and the descending branch of said first U-shaped conduit runs downward. from said second contact container. 3. A method according to either of claims 1 and 2, wherein isolated gaseous atmospheres are maintained in each of said contact vessels, and the fluidizing gases are supplied to the ascending branches of said U-shaped conduits to. from the atmospheres of said contact vessels respectively. <Desc / Clms Page number 32> 4. A method according to claims 1 to 3, wherein the portions of the U-shaped ducts below the point where the fluidizing gases are introduced into the upward branches of said ducts form safety branches which will allow sudden pressure fluctuations in the duct. either direction without tolerating mixing of the atmospheres in the containers. 5. A method according to claims 1 to 4, wherein a. Additional thinning agent is supplied directly to the bottom of the vessel receiving the upward transfer line where the density of the fluid is changed to control the flow rate of the solid, and in which the flow control is achieved by varying the proportion of the total supply of fluidizing agent which is supplied to the flow control line, compared to that which is supplied directly to the vessel. 6. A process according to claims 1 to 5, wherein pressure fluctuations in the system are minimized by introducing each fluidizing agent into the circulating fluidized solid stream in the same direction as the direction of flow. flow of the solid so as to avoid any tendency to reverse the normal direction of flow either of the solid or of the gas at the point of injection. 7. A method according to claims 1 to 6, for circulating a fluidized solid continuously in a closed circuit between two contact vessels containing fluid beds of solid and different gaseous atmospheres, which comprises removing of a dense column of solid fluidified downwards from the fluid bed of solid maintained in the first of said containers, the passage of the fluidized solid from the bottom of said dense column downwards to rise to the same level by a safety branch having the shape of a U-shaped siphon tube, the passage of the fluidified solid upwards from the ascending branch of said safety branch by an ascending transfer pipe leading to the fluid bed of solid maintained in the second container, adding sufficient aeration gas to the solid moving within said safety branch to maintain fluidity with maximum fluid density, introducing a fluidizing agent into the bottom of said ascending transfer pipe above the level of said safety branch so that the density of the fluidized solid in said ascending transfer pipe is appreciably less than that prevailing in the safety branch and less than that of the dense column of fluidized solid going therein up and down from the first-mentioned container, removing a second dense column of fluidized solid downwardly from the fluid bed in the second container and by a second U-shaped safety branch, to a second ascending transfer pipe leading from the bottom up in the first container, the introduction of a different fluidizing agent into the bottom of said second ascending transfer pipe above the level of said second branch safety so that the density of the fluidized solid which is therein is appreciably less than that prevailing in said dense second column, thereby establishing a pressure differential in each of said U-shaped tubes while maintaining the average density and the total fluidistatic pressure head of the fluidized solid in each of said dense columns, greater than the average density and the total fluidistatic pressure head of the fluidized solid in the respective connecting upward transfer conduits, which causes the solid to move downwards in high density columns and upwards in relatively low density transfer conduits, and to control the flow of circulating solid without the use of mechanical parts to regulate the flow of solid by varying the amount of fluidizing agent introduced into at least one of said upward transfer conduits to control the density of the solid therein. 8. A method according to claim 7, wherein the density of the fluidized solid in the safety branches and in the dense columns <Desc / Clms Page number 33> leading to it from top to bottom, is at least 25 pounds per cubic foot, and the fluidistatic pressure differential created in each of the safety branches below the point of injection of the additional thinning agent into the connecting riser and above the bottom of the safety branch, is at least 2 pounds per square inch. 9. A process according to claim 7, wherein the fluidified solid is a catalyst for the conversion of hydrocarbons, a process for the conversion of hydrocarbons is carried out in one of the vessels, the catalyst is regenerated in the other of the vessels. said vessels, and the level of the fluidized bed of catalyst in the regeneration vessel is maintained substantially constant by withdrawing dense fluidized catalyst from an upper portion of the bed into a supply pipe running up and down from said vessel. regeneration vessel, 10. A method according to claim 9, wherein the transfer line introducing catalyst into the regeneration vessel extends upwardly within said vessel and terminates at an intermediate level in the fluidized bed of catalyst therein. 11. An apparatus suitable for carrying out the method according to claims 1 to 10, comprising two contact receptacles arranged to contain fluid beds of solid, two open U-shaped conduits mounted below said receptacles and connecting each of the receptacles. to the other and arranged to contain a dense column of fluidized solid for the transfer of said solid from one of the containers to the other, each of said conduits comprising a straight vertical portion going downward from one of the containers by a curved bottom portion to a second straight vertical portion leading upward into the other container, conduits provided to introduce into said second vertical portion of each of said conduits a fluidizing agent to reduce the density of the fluidized solid which is therein and to pass upwardly with it into the contact container located above, a separate conduit for introducing an additional fluidizing agent directly into the bottom of one of said contact vessels and control means for varying the relative proportion of the total amount of fluidizing agent introduced directly into the bottom of said container and in the ascending duct which arrives there. from bottom to top, the entry point of each of said conduits for supplying the fluidizing agent into one or the other of the receptacles or the conduit which arrives there from bottom to top being situated at the bottom. above the curved bottom portion of the open conduit coming from the other container, whereby said curved bottom portions form, by transferring said dense columns of fluidized solid, U-shaped safety branches separating into- gaseous atmospheres in each of the contact vessels. 12. An apparatus according to claim 11, wherein a safety valve or tap is provided in the ascending branch of each of said U-shaped conduits below the point of introduction of the fluidifying agent with a device. control intended to completely shut off the flow of solid and gas in the event that the flow rate in the duct drops to an abnormally low rate. 13. An apparatus according to claim 11, wherein a purification or washing section is provided at the bottom of one of said contact receptacles, the section in the bottom of which penetrates the pipe descending from one of said conduits. U-shaped, and which is provided with means for injecting a purifying gas in order to wash the solids therein before passing through the descending branch of the U-shaped conduit. 14. An apparatus according to claims 11 and 12, wherein an extraction well is provided at the bottom of said second contact vessel, in which the solid is extracted and from which the solid passes into the descending branch of the other. U-shaped lead that gets there. 15. An apparatus for the catalytic conversion of hydrocarbons, <Desc / Clms Page number 34> comprising a cylindrical reaction vessel arranged vertically for containing a fluidized bed of finely divided solid catalyst, a scrub section disposed below said reaction vessel with means provided for introducing a scrub gas therein, a scrub section U-shaped open siphon tube extending downward from it and connecting the bottom of said scrubbing section to a vertical supply pipe disposed above said U-tube and leading downward up from the other end of it, a cylindrical regeneration vessel disposed vertically receiving the upper end of said pipe and intended to contain a fluidized bed of spent catalyst undergoing oxidative regeneration to burn off the combustible deposits formed on the catalyst during the conversion reaction, means for introducing a variable amount of fluidizing gas at the bottom of said vertical pipe above said U-shaped siphon tube so as to provide variable control of the density of the fluidized catalyst contained therein, separate means for introducing an oxidizing gas in the bottom of the regeneration vessel, an evacuation well to receive the catalyst overflowing from the fluid bed contained therein, a second vertical pipe receiving the catalyst coming from said evacuation well, a second open U-shaped siphon tube extending downwards from said pipe connecting the bottom of said vertical pipe to an upward transfer pipe leading to the reaction vessel, and means for introducing hydrocarbons intended to undergo cracking into the bottom of said ascending transfer pipe above the point where it meets said U-tube. 16. A combined process for the refining of petroleum oils, comprising distilling a crude petroleum oil into a series of fractions, subjecting one or more of said fractions to a conversion reaction, passing the residual distillation fraction of said crude oil into a fractionation tower and supplying each of said fractions or products of any conversion treatment thereof, preferably in the form of steam, to said fractionation tower at points below the entry point of said distillation residues, whereby said vapors wash said residue and separate streams of each desired product fraction are obtained. 17. A process according to claim 16, wherein one or more of the product streams from said fractionator are subjected to a conversion treatment and the products of said conversion treatments are supplied to said fractionation zone at a point. located below the entry point of said distillation residue. 18. A process according to claims 16 and 17, wherein the gas oil fraction coming from said fractionation zone is subjected to a catalytic cracking treatment, and the cracked products are supplied to said fractionation tower at a point located. below the point of entry of said distillation residue. 19. A process according to claim 18, wherein said fraction of gas oil is mixed with the fraction of virgin gas oil coming from the distillation phase of the crude product, before being subjected to the cracking treatment. 20. A process according to claims 18 and 19, wherein the cracking of the gas oil is carried out by the method and apparatus claimed in claims 1 to 15. 21. A process according to claims 16 to 20, wherein the naphtha fraction from the distillation of crude oil oil is subjected to a thermal or catalytic reforming treatment, and the reformed product is sent to said zone. fractionation. 22. A process according to any one of the preceding claims, in which the fraction of the bottoms from the fractionation zone is filtered and sent to storage. <Desc / Clms Page number 35> 23. A process according to claims 16 to 22, in which the naphtha fraction is subjected to a catalytic reforming treatment giving hydrogen, 24. A process according to claim 23, wherein the catalytic reforming treatment comprises the treatment of the naphtha fraction in the presence of a platinum group catalyst at a temperature between 800 and 1,000 F under a pressure between 400 and 1,000. pounds per square inch with a hydrogen supply of 30,000 to 12,000 standard cubic feet per barrel of naphtha. 25. A process according to claims 23 and 24, wherein the catalyst contains a catalytically active metal component comprising molybdenum, chromium, vanadium or tungsten. 26. A process according to claims 23 to 25, wherein the fraction of naphtha subjected to the catalytic reforming is a fraction of naphtha from the distillation phase or the fraction of naphtha from the fractionator of the product or a mixture of. some or all of these two fractions. 27. A method according to claim 26, wherein said naphtha fraction is a heavy naphtha fraction. 28. A process according to claims 23 to 27, wherein one or more of the other fractions obtained during the process from the distillation phase of the crude product or from the fractionation phase of product, are treated with the hydrogen obtained from said reforming phase, externally with respect to said distillation zone and to the product fractionation zone. 29. A process according to claim 28, wherein the treatment of the fraction with hydrogen is carried out by treating the fraction with the total effluent from the catalytic reforming phase. 30. A process according to claims 28 and 29, wherein the fraction treated with hydrogen is a light gas oil fraction or a heavy gas oil fraction withdrawn from the product fractionation zone. 31. A process according to claims 28 to 30, wherein the fraction of the reduced crude product bottoms from the distillation zone is treated with hydrogen to purify it and the washed reduced crude product is sent to the distillation zone. the split zone. 32. A process according to claim 31, wherein the fraction of reduced crude product is washed with the effluent from the reformer at the pressure prevailing in the reformer, and the pressure exerted on the fraction of washed crude product is reduced beforehand. passage through the product fractionation zone. 33. A process according to claims 31 and 32, wherein the hydrogen and stripped hydrocarbons removed from the reduced crude product are separated and the hydrogen is returned through the cycle to the catalytic reforming zone. 34. A process according to claims 28 to 33, wherein the gas oil fraction is treated with hydrogen before being subjected to the catalytic cracking treatment. 35. A process according to claims 28 to 34, wherein the gas oil fraction is subjected to catalytic cracking and the cracked products are treated with hydrogen. 36. A process according to claims 22 to 35, wherein a part of the crude oil fraction is subjected to a coking treatment including the production of heat, and the heat thus produced is supplied to the reforming zones. @ 37. A method according to any one of claims 23 to 36 wherein most of the heat required in the frac- <Desc / Clms Page number 36> Product fractionation is supplied in the form of sensible heat of the products of the various conversion phases which are fed there. 38. The process according to claims 16 to 37, wherein said product fractionation zone has a lower washing section and an upper combined absorption and fractionation section, said reduced raw product is supplied to the section. upper part of said washing section, said vapors of the different fractions and conversion products are brought to the lower part of said washing section, said distillate and said gas oil fractions are removed from said dry - Upper fractionation section, and said tails of reduced crude product are collected at the base of said lower washing section. 39. The method of claim 38, wherein a pressure is maintained higher at the top of said upper section than at the base of said lower section, 40. The process of claim 38, wherein said reduced crude product is washed by subjecting it to two phase heating, vaporization and washing action, first with vapors of the ef. - fluent from the catalytic cracking treatment of the gas oil fraction and secondly with the vapors of the other fractions and of the conversion products. 41. The process of claim 40, wherein the cracked gas oil fraction is washed in the product fractionation zone with vapors from the other fractions or conversion products which are fed there. 42. The process according to claim 41, wherein at least a portion of the cracked gas oil vapors is heavier but of lower temperature than said vapors of the other fractions or of the conversion products. 43. The process according to claims 16 to 42, wherein a fraction of the boiling point range of fuel oil obtained from said crude petroleum oil is subjected to heat treatment at high temperature under conditions. severe enough to cause at least a reduction in viscosity. 44. The process of claim 43, wherein said material is said tails of reduced raw product. 45. The method of claim 43, wherein the vapors produced during said heat treatment are sent to a tower portion of said fractionation zone where said washing takes place.