KR950008999B1 - Fcc process and apparatus having a low volume dilute phase disengagement zone in the reaction vessel - Google Patents

Abstract

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Description

Flow contact loading (FCC) method and apparatus with low volume dilution phase glass zone in reaction vessel

1 is a schematic diagram of a reactor / regenerator system for an FCC process arranged in accordance with the present invention.

* Explanation of symbols for main parts of the drawings

10: regeneration vessel 22, 26, 28: riser

24: raw material injection nozzle 34: reaction vessel

38, 88: collection chamber 42: cyclone separator

48, 86: deep rack 50: catalyst layer

52: stripping container 56, 58, 60, 62: baffle

70, 78: distributor 84: cyclone

92: collector

The present invention generally relates to a fluidized catalysic cracking (FCC) process of heavy hydrocarbon streams such as vacuum gas oil, reduced crude oil and the like. More specifically, the present invention relates to a method of reacting a hydrocarbon in an FCC reactor and separating the reaction product from the catalyst used in the reaction.

Fluid catalytic cracking of hydrocarbons is the main stay method for the production of gasoline and light hydrocarbon products from heavy hydrocarbon loadings such as vacuum gas oil or residual crude oil. High molecular weight hydrocarbons, such as heavy hydrocarbon raw materials, decompose and break off large hydrocarbon chains, producing lighter hydrocarbons. These light hydrocarbons are recovered as products and can be used immediately or further processed to increase the octane barrel yield compared to the filing hydrocarbon raw materials.

Basic installations or devices for the fluid catalytic cracking of hydrocarbons have been around since the early 1940s. The basic components of the FCC process include a reactor, a regenerator and a catalyst stripper. The reactor includes a contact zone in which the hydrocarbon raw material contacts the granular catalyst and a separation zone in which the product vapor in the decomposition reaction is separated from the catalyst. Further product separation takes place in the catalyst stripper, which receives the catalysts from the separation zone and removes non-activated hydrocarbons from these catalysts by countercurrent contact with steam or other stripping medium. The FCC process is carried out by contacting a starting material which is a source of vacuum gas oil, reducing stream or other relatively high boiling hydrocarbons with a catalyst of finely divided or particulate solid material. The catalyst is transported like a fluid by passing gas or vapor through the catalyst at a rate sufficient to produce the desired fluid transport system. Contact of the oil with the flowing material promotes the decomposition reaction. During the decomposition reaction, coke precipitates on the catalyst. Coke is made up of hydrogen and carbon and may contain other materials such as sulfur, metals, etc., with traces entering the process with the starting materials. The coke blocks the catalytic activity of the catalyst by blocking the active site on the catalyst where the separation reaction takes place. The catalyst is usually transferred from the stripper to the regenerator for the purpose of removing coke by an oxidation reaction using an oxygen containing gas. Catalysts having a reduced coke content relative to the catalyst in the stripper (hereinafter referred to as "ashing wood") are collected and returned to the reaction zone. Oxidation of coke from the catalyst surface releases a large amount of heat, some of which is discharged from the regenerator together with the gaseous product of coke oxidation, commonly referred to as flue gas.

The remaining heat leaves the regenerator with the regeneration catalyst. The fluidization catalyst circulates continuously from the reaction zone to the regeneration zone and then back to the reaction zone. Fluidization catalysts not only provide a catalytic function but also act as a medium for transferring heat to the zone. It can be said that the catalyst present in the reaction zone is consumed, that is to say that the catalyst is partially inactivated due to precipitation of coke on the catalyst. The details of the apparatus for transporting the catalyst between each of the contact zones, the regeneration zones and the stripping zones and between these different zones are well known to those skilled in the art.

A specific pressure balance is required between the two vessels to achieve catalyst flow between the reactor and the regenerator. The catalyst transfer from the reactor to the regenerator is controlled by the pressure in the upper zone of the reaction vessel and the catalyst head in the flow passage from the reaction vessel to the control valve which regulates the amount of catalyst entering the regenerator. Similarly, catalyst transfer from the regenerator to the riser is controlled by the catalyst head in the input in the regeneration vessel and in the flow passage from the regeneration vessel to the bottom of the riser. The pressure difference between the bottom of the riser and the top of the reaction vessel is usually 27.5-68.9 kPa (4-10 psi). As the pressure in the regenerator increases, a corresponding pressure increase also occurs in the reactor. The increase in pressure of the regenerator to the reactor is limited by the catalyst head in the catalyst flow path from the reactor to the regenerator. Pressure interdependence sometimes does not provide the most favorable pressure conditions for both the reactor and the regenerator. Low pressure conditions are generally advantageous in the riser. Low pressure in the riser improves the evaporation of hydrocarbons and the product yield. High pressure conditions, on the other hand, are advantageous in regenerators, increasing coke burning capacity and improving kinetics for coke burning. However, it is not possible to increase the regenerator / reactor pressure differential beyond a certain limit due to the altitude of the solute tip that can be obtained in the reaction vessel and the stripper.

The rate of conversion of the feedstock within the reaction zone is controlled by adjusting the temperature of the catalyst, the activity of the catalyst, the amount of catalyst (ie catalyst to oil ratio) and the connection time between the catalyst and the raw material. The most common method of adjusting the reaction temperature is to adjust the rate of catalyst circulation from the regeneration zone to the reaction zone, which also causes a change in the catalyst to oil ratio at the same time as the reaction temperature changes. Therefore, when it is desired to increase the conversion rate, it is effective to increase the flow rate of the fluid catalyst circulating from the regenerator to the reactor.

The hydrocarbon product of the FCC reaction is recovered in vapor form and sent to a product recovery facility. Such equipment typically includes a main column that cools hydrocarbon vapors from the reactor and recovers a series of heavy cracked products, typically containing bottoms, circulating oil and heavy gasoline. Lighter material from the main column enters the flocculation compartment and is again separated into additional product streams.

Reduction of product loss and improvement in control of regeneration temperature have been achieved by providing multistage catalyst stripping and by raising the temperature of the catalyst particles to strip from the product and other combustible compounds. Both of these methods increase the amount of low molecular weight product stripped from the catalyst and reduce the amount of combustibles in the furnace regenerator. There are several known multistage process arrangements for stripping and heating the spent catalyst to raise the temperature of the stripping zone. A method of raising the temperature of the stripping zone by mixing the spent catalyst with the hot regenerated catalyst from the regeneration zone is known. It is being proposed more and more frequently.

As the FCC unit developed, the internal temperature of the reaction zone gradually increased. The diameter is usually used at a temperature of 525 ° C (975 ° F). At higher temperatures, the gasoline component is generally lost as the gasoline component decomposes into a lighter component by both the contact mechanism and the correct thermal mechanism. At 525 ° C. (975 ° F.), possible gasoline yields are generally reduced by 1% due to the pyrolysis of gasoline components and conversion to lighter hydrocarbon gases. As temperature increases, for example, at 550 ° C. (1025 ° F.), most feedstocks lose up to 6% more gasoline yield due to pyrolysis of gasoline components.

To reduce the loss of product by pyrolysis, one improvement to the FCC unit is to use cracking in the riser. In riser cracking, the regenerated catalyst and catalytic material enter the piped reactor and are conveyed upward by the expansion of the gas due to the evaporation of hydrocarbons and other flow media (if present) during contact with the hot catalyst. The riser cracks provide good initial contact of the catalyst with the oil, and also allow the contact time between the catalyst and the oil to be more strictly zero by eliminating turbulence and backmixing that can change the catalyst residence time. Today's average riser crack zone has a catalyst-oil contact time of 1 to 5 seconds. Many riser devices use lift gas as an aid to provide a uniform catalyst flow. Lift gas is used to accelerate the catalyst in the initial zone of the riser prior to the introduction of the feed, thereby reducing the turbulence that can vary the contact time between the catalyst and the hydrocarbon.

The advantage of using lift gas to pre-accelerate and regulate the regeneration catalyst in the riser type transition zone is known. Lift gases typically have low concentrations of heavy hydrocarbons. In other words, it is preferable to avoid hydrocarbons having a molecular weight of C ₃ or higher as the lift gas. In particular, highly reactive types of moles such as C ₃ + C olefins are unsuitable as lift gases. Therefore, lift gas streams consisting of steam and light hydrocarbons are generally used.

Regardless of the use of lift gas, ascending tube decomposition has provided substantial benefits to the operation of the FCC unit. In summary, the short contact time in the reaction riser controls the degree of decomposition occurring in the riser and improves the mixing to provide a more homogeneous mixture of catalyst and raw material. A more complete distribution prevents the contact time difference between the catalyst and the raw material in which part of the raw material contacts the catalyst for a longer time than other parts of the raw material, on the cross section of the riser. Both the short contact time between the feedstock and the catalyst and the more uniform average contact time for all feedstocks allowed the overdegradation to be controlled or eliminated in the reaction riser.

Unfortunately, much of what can be achieved in the reaction riser with respect to the uniformity of the raw material contact and the controlled contact time can be lost when the catalyst is separated from the hydrocarbon vapor. Catalysts and hydrocarbons must be separated as they exit the riser. Also in the conventional riser decomposition operation, the product in the riser was discharged into a large container. This vessel acts as a glass chamber and is still called a reaction vessel despite most of the reaction taking place in the reaction riser. This reaction vessel has a large volume. The steam entering the reaction vessel mixes well in large volumes and therefore has a wide residence time distribution, whereby a significant portion of the product fraction can have a relatively long residence time. Product fractions with an extended residence time can again undergo contact and thermal decomposition to produce unwanted low molecular weight materials.

In order to further control the contact time between the catalyst and the raw material vapor, research has been conducted using a cyclone directly connected to the end of the reaction riser (see US Patent No. 4,737,346). The direct connection of the cyclone to the riser thus quickly separates most of the product vapor from the catalyst. Therefore, it is possible to precisely control the contact time for most of the raw material vapors. One problem with using cyclones directly at the outlet of the reaction riser is the need for a system that can handle pressure surges from the riser (see US Pat. No. 4,624,772). This pressure rise and consequently an excessive increase in the catalyst load inside the cyclone overloads the cyclone such that an unacceptable amount of particulate catalyst is carried along with the reactor vapor to the downstream separation facility. Therefore, there is a way to considerably entangle the arrangement and arrangement for a direct-type cyclone and provide a process arrangement where a significant amount of reactor vapor can enter the open space of the reaction vessel or a temporary catalyst overload in the cyclone system. A number of device arrangement methods have been proposed that make it possible to operate the cyclone system satisfactorily.

Direct cyclone systems can help control the contact time of the catalyst and feed steam, but do not completely remove hydrocarbon vapors present in the open space of the reaction vessel. The product vapor is still present in this open space as stripped hydrocarbon vapor form which is removed from the catalyst and passes upward in the open space above the stripping zone. In addition, the amount of hydrocarbon vapor is increased by the direct-attached cyclone apparatus to allow the feed vapor to enter the open space containing the cyclone. The dilution volume of the reaction vessel remains unchanged when the tandem cyclone is used, and because less hydrocarbon vapor enters the dilution phase volume from the riser, the hydrocarbon vapor entering the dilution phase volume is When used they will be in dilution phase volume for much longer [the terms "dense phase" and "dilution" catalyst as used herein mean catalyst density in a particular zone. The term "dilute phase" means a catalyst density of less than ^{320㎏ / ㎥ (20lbs / ft 3} ), the term "dense phase" means a catalyst density of ^{480㎏ / ㎥ (30lbs / ft 3} ). If the catalyst density is 320 to 480 kg / m ³ (20-30 lbs / ft ³ ), it should be determined whether it is a dense or dilute phase according to the density of the catalyst in the adjacent zone or zone. In other words, when a tandem cyclone system is used, less product vapor enters the open space of the reaction vessel, but the introduced vapor stays in the reaction vessel for much longer. As a result, any raw material remaining in the reaction vessel is overdegraded and virtually lost.

In addition, a very slow gas flow rate through the reaction vessel promotes precipitation of coke on the interior of the vessel. The long residence time of the relatively hot heavy hydrocarbons in the upper compartment of the reaction vessel promotes the formation of coke. Such coke precipitation interferes with the function of the reaction vessel by thickening the precipitate inside the vessel to insulate a portion of the metal shell and locally cool it. Such local cooling sites can promote coagulation of corrosive materials and damage the reaction vessel. In addition, large coke precipitates can often be broken into large chunks, creating another problem that these pieces can block the flow of catalyst through the vessel or conduit.

A device known to promote rapid separation between catalyst and vapor in a reaction vessel is a ballistic separation device called a vented riser (see US Pat. No. 4,792,437). The basic form of the vented riser structure consists of a vertical conduit portion at the end of the riser and an opening pointing up towards the reactor provided with a plurality of cyclone inlets surrounding the outside of the riser near the opening end. The apparatus works by blowing high momentum catalyst particles above the open end of the riser where gas collection takes place. Rapid separation between gas and vapor occurs due to the relatively low density of the gas, which can change direction quickly and enter the inlet near the riser (see US Pat. No. 4,295,961), while heavier catalyst particles Continues along the straight trajectory formed by the vertical portion of the rising conduit. The vented riser provides the advantage of providing quick separation between the catalyst and the vapor, while eliminating all of the useless areas in the reaction vessel where coke can form. However, the vented riser still has the drawback of operating within a large open space of the reaction vessel.

The problem discussed by the present invention is the need to expedite the separation between the product vapor and the catalyst leaving the FCC riser in a system that minimizes overdegradation of the vapor phase product and minimizes the entrainment of the vapor phase product and particulate catalyst.

It is an object of the present invention to reduce the hydrocarbon residence time in the reaction vessel.

Another object of the present invention is to improve the function of the vent type riser in the separation of the catalyst and hydrocarbon vapor.

Another object of the invention is to control the residence time in the reaction vessel zone of the FCC reaction zone.

The present invention provides an FCC arrangement in which the outlet end of the reaction riser is discharged into the upper portion of the reaction vessel, which is possible as an open vitrification zone, and the discharge end of the reaction riser is positioned close to the top of the dense catalyst in the reaction vessel. It is about a method. This arrangement is facilitated by installing cyclone or gas solid separation outside the reaction vessel. It has been found that when the belt riser is used with an externally located cyclone or separation device, the top level of the dense catalyst bed can be kept close to the top of the reaction vessel. The high level of the dense catalyst bed relative to the riser outlet reduces the dilution phase volume in the vitrification vessel, thereby reducing the hydrocarbon residence time and reducing the required reaction vessel height above the dense bed level.

Many advantages can be obtained by reducing the height of the dilution phase in the reaction vessel. Reduced dilution phase height can shorten the reaction vessel and its overall height. In addition, additional heights can be used to maintain longer vertical lengths for the dense catalyst bed while maintaining the overall height of the reaction vessel. Further stripping of the dense catalyst bed can further carry out catalyst stripping. In addition, the higher the dense catalyst height, the higher the pressure drop between the reaction vessel and the regenerator control valve, thereby maintaining a higher regenerator pressure without raising the reactor zone. In addition, it is possible to control the total residence time of the hydrocarbon vapor in the reaction vessel because it can make a significant change in the level of the dense catalyst bed. This makes it possible to use very short residence times for certain raw materials and to increase residence times for more refractory raw materials without changing the feed rate to the reaction riser. This advantage shows that the present invention can provide almost the same improvement as that provided by the tandem cyclone in reducing overdegradation while providing tremendous flexibility with increasing reliability of the open-ended reaction riser. .

Therefore, in one embodiment, the present invention relates to a process for fluid catalytic cracking of FCC materials of the present invention. The method includes feeding FCC feedstock and catalyst particles to a reaction riser having an effective riser outlet diameter, passing the catalyst and feedstock upwards through the riser to convert the feedstock into a gaseous product stream, and by adhering coke. Producing spent catalyst particles. The mixture of spent catalyst particles and gaseous product is discharged upward into the reaction vessel upwards from the top of the reaction vessel to the discharge end of the riser located below 1-12 times the diameter of the riser outlet and consumed the initial beak of the spent catalyst from the gaseous product. To provide. The separated catalyst passes downward through the reaction vessel, is maintained as a dense catalyst layer in the reaction vessel to form a stripping zone, and the stripping gas passes upward through the reaction vessel. The upper surface of the catalyst bed is between a minimum spacing of 0.6 m (2 ft) at the riser outlet half or one time below the riser diameter and a maximum spacing of 2.4 m (8 ft) at the riser outlet end to four times below the riser diameter. Is maintained at. The mixture of gaseous product, stripping fluid and spent catalyst particles is recovered from the reaction vessel and sent to a particle separator located outside of the reaction vessel, where the gaseous components are separated from the spent catalyst which is eventually returned to the regeneration zone. The spent catalyst particles are fed from the stripping zone to the regeneration zone and contacted with the regeneration gas in the regeneration zone to combust the coke from the catalyst particles and produce regenerated catalyst particles to be sent to the reaction riser.

In addition, the present invention may be described in connection with an apparatus. The device is an upwardly rising conduit with an upwardly open end having an outlet opening corresponding to a diameter equivalent to that disclosed herein as "elevating pipe exit diameter", which surrounds the exit end and rises at the outlet end of the rising pipe. A reaction vessel having an upper end located 5-12 times above the diameter of the tube outlet. Gas solid separation devices are located outside the reaction vessel, and these devices have an inlet, a gas outlet and a solid outlet. Means are provided for collecting the mixture of gas and catalyst in the upper half of the reaction vessel and directing the mixture of catalyst and gas to the inlet of the separation apparatus. Means are provided for conveying catalyst particles from the collector to the reaction vessel, and means are provided for recovering the catalyst from the bottom of the reaction vessel and transferring the catalyst to the regeneration vessel.

The present invention generally relates to the reactor side of an FCC process. The present invention is applicable to most FCC methods used for the separation of hard or heavy FCC raw materials. The method and apparatus of the present invention may be used to modify the operation and placement of existing FCC units or may be used in the design of newly assembled FCC units.

The present invention uses the same general elements as more FCC units. The reaction riser provides the first reaction zone. The reaction vessel and cyclone separator remove catalyst particles from the gaseous product vapors. The stripping zone removes most of the adsorbed catalyst particles from the surface of the catalyst. The spent catalyst from the stripping zone is regenerated in a regeneration zone having at least one regeneration step. Regenerated catalyst in the regeneration zone is used in the reaction riser. Many different arrangements may be used for the elements of the reactor and regenerator zones. The description of specific reactors and regenerators disclosed herein is not intended to limit the invention, except as specifically disclosed in the claims.

An overview of the basic process operations will be understood with reference to the drawings. The regeneration catalyst is conveyed from the lower portion 12 of the regeneration vessel 10 through the conduit 14 to the Y-shaped zone 18 at a controlled rate by the control valve 16. Lift gas injected through conduit 20 to the bottom of Y-shaped zone 18 carries catalyst up through lower riser zone 22. The raw material is injected into the riser above the lower riser zone 22 by the raw material injection nozzle 24.

The mixture of raw material, catalyst and lift gas passes through the middle section 26 of the riser and ends at the upper inner riser section 30 ending in the upward outlet end 30 located in the dilution phase region 32 of the reaction vessel 34 ( 28). The gas and catalyst are separated in the dilution phase zone 32. Conduit 36 collects gas from the dilution phase zone and transfers it to collection chamber 38. From the collection chamber 38, the T-pipe apparatus 40 distributes the gas still containing a small amount of catalyst particles to the pair of cyclone separators 42. Relatively pure product vapor is recovered by the manifold 44 from the outlet of the cyclone 42 and recovered from the process via the outlet 46. The catalyst separated by the cyclone separator 42 is conveyed back to the reaction vessel 34 by a dip pipe conduit 48. The spent catalyst from the dilution phase zone 32 and the dippipe conduit forms a dense catalyst layer 50 at the bottom of the reaction vessel 34. The separated catalyst passes down the vessel at an average rate of 98 kg / m 2 / sec (20 lb / ft ² / sec). The dense catalyst bed extends downward into stripping vessel 52 which acts as a stripping zone. The stripping fluid enters the bottom of the stripping vessel 52 through the distributor 54 and moves up the stripping vessel and the reaction vessel with countercurrent flow to the downwardly moving catalyst. As the catalyst moves down, it passes over reactor stripping baffles 56 and 58 and stripper baffles 60 and 62 and at a rate regulated by control valve 66 via conduit 64. Conveyed into the upper region 68 of 10. The catalyst particles are in contact with the oxygen containing gas in the upper zone 68 of the regeneration zone. The separator 70 receives the oxygen-containing gas from the conduit 71 and distributes the gas to the entire cross section of the regeneration vessel. The regenerated catalyst is recovered from the upper zone 68 and conveyed through the conduit 72 to the lower portion 12 of the regeneration vessel at a controlled rate by the control valve 74.

At the lower portion 12, the catalyst enters the vessel via conduit 76 and is contacted with additional regeneration gas dispersed throughout the cross section of the vessel by the domed distributor 78. In the lower compartment 12 the catalyst is completely regenerated and recovered by the conduit 14 according to the method described above. In the bottom regeneration zone 12 the coke combustion product rises and enters the top regenerator zone 68 via a series of internal flues 80. Flue gas from the lower zone 12 is mixed with flue gas generated in the upper zone 68 and recovered through the inlet 82 of the cyclone 84. The flue gas entering cyclone 84 contains a small amount of catalyst particulate, which is removed by the cyclone and returned to the regenerator by dip leg 86. Flue gas exiting the cyclone is collected in collection chamber 88 and leaves the regenerator through conduit 90.

The reaction riser of the present invention is installed to effect initial separation between the catalyst and gas phase components in the riser. The term "gas phase component" includes lift gas, product gas and steam as well as unconverted raw material components. 1 shows the invention using a phase tube apparatus with a lift gas zone 22. Lift gas zones need not be provided in the riser to enjoy the advantages of the present invention. However, the end of the riser should end with one or more upward openings which discharge the catalyst and gaseous mixture upwards into the dilution phase zone of the reaction vessel. The riser opening end is the opening end of a normal vented riser as disclosed in the prior art patent of this application or in any other arrangement of openings that provides substantial catalytic separation from the gaseous material in the dilution phase zone of the reaction vessel. It may be an end. It is important to drain the catalyst upwards to minimize the gap between the outlet end of the riser and the top of the dense catalyst bed in the reaction vessel. The flow regime inside the riser affects the separation at the ends of the riser. Typically, the rate of catalyst circulation through the riser, the input of feedstock and any lift gas entering the riser, is about 48-320 kg / m ³ (3-20 lbs / ft ³ ) for the catalyst and gas phase mixture and about It provides an average speed of 3-30 m / sec (10-100 ft / sec). The length of the riser is set to provide a residence time of 0.5 to 5 seconds under these average flow rate conditions.

The rate at which the catalyst and gaseous mixture exits the end 30 of the riser affects the placement of the end of the riser relative to the top of the reaction vessel. This distance, indicated by the letter "A" in FIG. 1, is provided based on the flow rate to the riser. The interval "A" should be as short as possible in order to minimize the dilution volume of the catalyst in the reaction vessel. Nevertheless, avoid the direct impact and corrosion of the top of the reaction vessel, and the reconstruction of the gas stream collected from the upper zone of the reaction vessel with catalyst particles where the discharge of catalyst from the end of the riser is separated by the initial discharge from the riser. Some space is required between the ends of the riser to prevent entrainment and to provide separation. Since the reaction riser is typically designed for a narrow discharge speed range of 6.1 to 30.5 m / s (20-100 ft / s), the spacing “A” can be determined based on the riser outlet diameter. In order to avoid corrosion of the upper surface of the reaction vessel and to facilitate the initial separation of the catalyst from the gaseous components, the spacing "A" should be at least one time the riser outlet diameter. Avoidance of re-entry of the catalyst after discharge of the riser depends both on the riser speed and the flow density of the catalyst as it passes down the reaction vessel. For most practical catalyst density ranges in the riser, riser outlet diameters range from 1.5 to 12 times, preferably 1.5 to 6 times the riser diameter, from about 244 to 976 kg / m 2 / sec (50-200 lb / ft). Suitable for the "A" dimension for the flow catalyst density (in some cases referred to as "catalyst flux") of ² / sec).

The total volume per dilution in the reaction vessel is the diameter of the reaction vessel, the distance from the end of the riser to the top of the reaction vessel, i.e. the "A" dimension and the distance from the discharge end of the riser to the top of the dense bed level in the reaction vessel ( As indicated by the dimension "B" in the figure). In order to prevent the catalyst particles from re-blowing back to the gas recovered from the reaction vessel, a minimum gap is required between the top of the reaction riser and the top of the catalyst bed level. This dimension mainly depends on the apparent speed of the gas flowing upward through the dense layer 50. In order to minimize the possibility of entraining the gaseous compound through the dense bed 50, the apparent speed is generally less than 0.15 m / sec (0.5 ft / sec). The gaseous components passing upward through the dense bed 50 consist of stripping fluid and hydrocarbons desorbed from the catalyst surface. The amount of stripping gas entering the stripping vessel is typically proportional to the void volume in the catalyst. For most rational catalyst to oil ratios in the riser, the amount of stripping gas that must be added to replace the pore volume of the catalyst does not exceed 6% by weight of the raw material amount. Therefore, the relative ratio of the catalyst and stripping fluid as well as the stripped fluid as it passes through the reaction vessel upside down remains relatively constant. Thus, the main variable in controlling the apparent gas velocity passing upward through the dense catalyst bed is the diameter of the reaction vessel. 0.6, as long as the apparent velocity of the gas rising through the dense bed 50 is maintained in the range of less than 0.15 m / sec (0.5 ft / sec), preferably less than about 0.03 m / sec (0.1 ft / sec) A spacing "B" of -2.4 m (2-8 ft) prevents entrainment of gas and catalyst back out of the reaction vessel. For most reaction risers, this 0.6-2.4 m (2-8 ft) corresponds to about one to four times the diameter of the riser outlet.

In addition, the manner in which the gaseous vapor is recovered from the dilution volume of the reaction vessel also affects the initial separation and the degree of entrainment of the re-compound obtained in the reaction vessel. In order to improve this vitrification and to prevent entrainment, Figure 1 suggests the use of an annular collector 92 surrounding the end 30 of the riser. Collector 92 is supported by recovery conduit 36 from reaction vessel top 34. The recovery conduit 36 is symmetrically spaced about the annular collector and communicates with the annular collector through a plurality of symmetrically spaced openings to recover gaseous components evenly throughout the perimeter of the reaction riser. Both the gaseous components from the stripping gas and the reaction riser are recovered by the annular collector 92 according to the process arrangement shown in FIG. Product gas from conduit 36 is all sent to cyclone 42.

FIG. 1 shows a process arrangement for transferring gas from conduit 36 to a cyclone that prevents unbalanced distribution of catalyst and gas mixtures for different cyclones. The simplest way to connect a gas conduit with a cyclone is to connect one conduit directly to the corresponding cyclone. This arrangement also has the advantage of minimizing the flow path between the annular collector of the riser and the cyclone in which the final separation of catalyst and gas takes place. However, although no reason was found, it was found that a mixture of catalyst and gas taken from the reactor through a series of conduits could optionally flow into one conduit. As a result, heavier loads of catalyst and gas can overload the cyclone through which the flow runs. For this reason, FIG. 1 illustrates the use of a collection chamber 38 that collectively collects gas from all cyclone conduits 36 and redistributes this gas to individual cyclones. Although the use of the collection chamber 38 and the T-type zone 40 increases the residence time when the catalyst and gas mixture flows from the reaction vessel to the cyclone inlet, this slight increase in residence time recovers from the cyclone. It has no real impact on quality. In addition, the prevention of unbalanced distribution can also be achieved by the use of catalysts and gas separation devices in addition to cyclones.

The catalyst recovered by the cyclone can be returned to the process at any convenient location. Regardless of which type of gas and catalyst separation device is used, the catalyst separated therefrom is returned to the process. The catalyst can be returned to any point in the process that puts it into the catalyst circulation path. The figure shows the use of a conventional cyclone with a dip rack 48 conveying near the upper layer water level 51 of the dense layer region 50. The catalyst is returned to the dense layer in the reaction vessel or stripping.

Once the cyclone is removed from the reaction vessel, the diameter of the reaction vessel is no longer affected by the need to provide adequate space for the separation device therein. Thus, the diameter of the reaction vessel may be determined based on the apparent velocity of the gas passing upward through the dense bed and the catalyst flux entering the reaction vessel. As mentioned above, the criteria for both of these parameters allow for a smaller reactor diameter than was used in the prior art. Smaller reaction vessel diameters further reduce the volume of the dilution phase in the reaction vessel. When the present invention is used with a new reaction vessel, its diameter can be kept low such that the average residence time in the dilution phase of the reaction vessel is less than 3 seconds. In addition, since the apparent speed and catalyst flux are affected by well-known catalyst density ranges and speed conditions in the riser, the diameter of the reaction vessel initially designed according to the present invention is 3 to 5 of the outlet diameter of the riser. It is preferable to be between ships. Alternatively, the dilution phase volume of the reaction vessel above the top of the catalyst bed can be maintained at less than five times the reaction riser volume through which the raw material passes.

The catalyst, initially separated from the gaseous constituent, passes into the vessel downwards as it enters the reaction vessel as described above. The catalyst is in contact with a series of baffles that enhance the contact of the catalyst with the stripping gas passing through the vessel and preferably upwardly through the vessel. In the embodiment of the invention shown in FIG. 1, the catalyst passes through a stripping zone at the top of the vessel and a separate stripping vessel located below it. 1 shows baffles 56 and 62 located outside of the vessel wall and baffles 58 and 60 located longitudinally downward of the riser along the lower portion of the reaction vessel and stripping vessel. These stripping baffles typically act to drop the catalyst from side to side as the catalyst passes through the vessel and increase the contact of the catalyst particles as the stripping vapor passes upwards in countercurrent contact with the catalyst. The dense layer 50 has a relatively long length in the reaction vessel 34. There is no requirement for a long dense bed length in the reaction vessel, and the dense bed length shown in FIG. 1 derives from the type of device shown in FIG. 1 depicts a retrofit of an existing regenerator and reactor zone in accordance with the present invention, wherein the tangent length of the reaction vessel is defined by a conventional process arrangement in which the cyclone separator is placed inside the reaction vessel. When the method of the present invention is used in the initial design of the reaction vessel, the tangent length can be substantially reduced such that the upper dense bed level 51 is close to the top of the stripper vessel.

However, the thickness of the dense catalyst layer in the reactor 34 increases the overall height of the dense catalyst on the control valve 66. The additional thickness of this dense bed catalyst can be advantageously used in FCC processes. First, the additional length of dense bed catalyst provides an extended zone that increases the contact between the stripping fluid and the catalyst. Therefore, more stripping can be obtained by the extended length of the dense catalyst layer. In addition, the hydrostatic catalyst head from the top surface 51 to the control valve 66 causes a relatively large pressure drop between the control valve 66 and the dense bed level 51. This pressure drop may be more than 48 kPa (7 psi) total. This additional pressure allows the regenerator to be operated at higher pressure than the reactor zone. As mentioned above, operating the regeneration zone at higher pressures and operating the reaction zone at lower pressures has significant advantages.

In addition, an additional thickness of dense layer can be used to introduce the hot stripping zone. This hot stripping zone can utilize the catalyst from the regeneration zone to supply heat to the stripping zone and increase the desorption of hydrocarbons and volatile components from the surface of the catalyst. Suitable lift systems can be used to transport the catalyst from the regeneration zone to the desired highly upstream stripping zone.

The catalyst is recovered in the stripping zone and sent to the regeneration zone. Any known regenerator apparatus for removing coke from the catalyst particles by oxidative combustion of the coke and returning the catalyst particles to the reaction riser can be used.

Claims (11)

(a) feeding the feedstock for catalytic catalytic cracking (FCC) and catalyst particles into a reaction riser having a riser outlet diameter, and transporting the catalyst and feedstock through the riser to convert the feedstock into a gaseous product and Producing spent catalyst particles by coke deposition on the catalyst particles; (b) spent catalyst into the rise reaction vessel from the discharge end of the riser located 5-12 times below the riser outlet diameter at the top of the reaction vessel; Evacuating the primary mixture of particles and gaseous product upwards to provide initial separation of spent catalyst from the gaseous product, (c) passing the separated catalyst downward through the reaction vessel, and (d) The catalyst is held in the reaction vessel as a dense catalyst layer to form a stripping zone, and the stripping gas passes upward through the stripping zone. (E) maintaining the top surface of the dense catalyst bed at a distance of 1 to 4 times the diameter of the riser outlet at the riser outlet, and (f) regenerating the catalyst stripped from the stripping zone. Feeding the spent catalyst with the regeneration gas in the regeneration zone to combust coke from the catalyst particles to produce regenerated catalyst particles to be sent to the reaction riser, and (g) gaseous product, stripping Recovering a secondary mixture of fluid and spent catalyst particles from the reaction vessel, transferring the mixture to a particle separator located outside of the reaction vessel, and separating gaseous components from the spent catalyst particles in the separator zone. Method of fluid contact hazard of FCC raw material consisting of. The method of claim 1 wherein the primary mixture is discharged from the riser at a speed ranging from 6.1 to 30.5 m / s (20-100 ft / s). The method of claim 1, wherein the secondary gas mixture is withdrawn from an annular collector surrounding the riser end and the particle separator comprises one or more cyclones. The method of claim 1 wherein the product withdrawn in step (b) has an average residence time of less than 3 seconds in the reaction vessel. The method of claim 1 wherein the reaction vessel has a diameter three to five times the diameter of the outlet of the riser. The method of claim 1 wherein the pressure drop from the bottom of the reaction vessel to the top of the dense layer is at least 48.3 kPa (7 psi). The process of claim 1 wherein the separated catalyst passes downward through the reaction vessel at an average speed of less than 98 kg / m 2 / sec (20 lb / ft ² / sec), and the stripping gas passes through the reaction vessel at 0.15 m / sec (0.5 ft). Per second) passing upward at an average apparent speed of less than. (a) an upward ascending conduit having an upward outlet end with a uniform outlet diameter, (b) a reaction enclosing the outlet end, the upper end of which is located 1 to 12 times above the riser outlet diameter at the outlet end; A vessel, (c) a gas solids separation device external to the reaction vessel and having an inlet, a gas outlet and a solid outlet, (d) collecting a mixture of gas and catalyst from the site of the reaction vessel, the catalyst and gas Means for delivering the mixture to the inlet of the separation device, (e) means for introducing a gaseous stripping medium into the reaction vessel below the outlet end, and (f) catalyst particles from the collector and from the bottom of the glass vessel to the reaction vessel. FCC apparatus of the raw material for fluid catalytic cracking (FCC) which consists of a means for conveying. 9. The annular collector of claim 8, wherein the gas solids separator comprises one or more cyclone separators located outside of the reaction vessel, and wherein the means for collecting the mixture of gas and catalyst surrounds the outlet end of the riser. Apparatus comprising a. 9. The apparatus of claim 8, wherein a dip leg transfers catalyst from each catalyst outlet to the reaction vessel, the outlet of each dipleg being located below the riser outlet. 2. The process of claim 1 wherein the dilution phase volume of the glass zone above the top of the catalyst is less than five times the volume of the reaction riser between the point at which raw material enters the riser and the discharge end.