JPH055875B2 - - Google Patents

Description

[Detailed description of the invention]

FIELD OF THE INVENTION This invention relates to the production of olefins and aromatic hydrocarbons from hydrocarbon feedstocks. In particular, the present invention relates to the production of olefins and aromatic hydrocarbons by catalytic cracking alone or catalytic cracking of hydrocarbons followed by dehydrogenation. Furthermore, the present invention provides for preferentially producing olefins having 3 or more carbon atoms and/or for producing aromatic hydrocarbons, in particular benzene, in the presence of an entrained flow of a thermally supported solid catalyst. This invention relates to a method for thermally decomposing hydrocarbons in a short residence time. BACKGROUND OF THE INVENTION It is known that natural hydrocarbons are decomposed at high temperatures to produce valuable olefinic materials such as ethylene and propylene. The increased demand for propylene-based plastics compared to ethylene-based plastics means that when hydrocarbons are broken down into olefins,
Beneficially acting to improve propylene yield. Additionally, higher olefins, such as C4 olefins, are important precursors for providing high octane blending components. That is, C4 olefin is MTBE
is a precursor for the production and alkylation of However, when heavy hydrocarbon feedstocks are cracked non-catalytically to olefins, it is virtually impossible to achieve the desired co-product ratio to meet demand.
That is, the yield of propylene to ethylene is rarely greater than 0.55. High yields are achievable only in low hydrocarbon conversions that increase regeneration costs and reduce supply. One well-known non-catalytic cracking method is pyrolysis, which generally occurs at high temperatures in the presence of water vapor. The mechanism for obtaining olefins by pyrolysis is explained in terms of free radicals. At high temperatures,
Radical initiation occurs by homolysis of carbon-carbon bonds. Once started, the free radicals are 2
experience two major reactions. That is, the first is the fragmentation of the radical at the beta position, and the second is the termination of the reaction by abstracting hydrogen. Cleavage at the beta position will continue to the point where methyl groups are generated 90% of the time. Methyl groups will extract hydrogen atoms from other molecules, producing methane and other free radicals. Ethylene and methane are typical products from such free radical pyrolysis reactions. Long-chain radicals extract hydrogen from molecules about 10% of the time.
Produces C3-C7 paraffins and olefins. Pyrolysis therefore produces ethylene in high yields compared to the higher olefins produced primarily as a result of branched hydrocarbons in the initial hydrocarbon feedstock. In order to increase the production of C3 and higher olefins, light hydrocarbons containing at least one alkane have been pyrolyzed in the presence of solid catalysts containing hydrogen sulfide and silica. [Klolts, U.S. Patent No.
4,471,151]. The contact material employed, for example silica gel, preferably has a large surface area, ie at least 50 m ² /gm. H2S concentrations of 0.1 to 10 mole percent relative to the alkane feedstock are generally employed in the process. It is speculated in the US patent that the increased cracking is due to the high surface material acting as a catalyst to decompose H2S. As a result, conversion levels are increased and improved selectivity for the desired product is obtained. However, improved selectivity to propylene was demonstrated only when cracking n-butane. The solid contact materials employed in said US patent are suitable only for fixed bed operations, but are not suitable for fluidized bed operations due to their extremely low mechanical stability. Therefore, the solid catalyst of said US patent continues to have the disadvantages of typical catalytic dehydrogenation catalysts designed for fixed beds.
These dehydrogenation catalysts are large size, diffusion limited catalysts and cannot be continuously regenerated in a circulating wheel system. Fluidized catalytic cracking (FCC) units may also be employed to catalytically produce C3 and higher compounds. FCC devices use acidic decomposition catalysts to increase the production of C3-C7 compounds based on a carbonium ion mechanism compared to a free radical pyrolysis reaction mechanism. However, the acidic decomposition activity of the catalyst, in addition to increasing decomposition and isomerization, promotes rapid hydrogen transfer and
As a result, the yield of paraffin is higher than that of olefin. However, the properties of the catalytic cracking unit itself are suitable for paraffin production. The typical definition of residence time in a catalytic cracking operation is the time that the feedstock is in contact with the catalyst itself. This definition is acceptable when the temperature is so low that no appreciable thermal reaction occurs. However, the thermal and catalytic reactions proceed in parallel. While catalyst separation terminates the catalytic part of the reaction, the thermal reaction (pyrolysis) continues until the temperature drops to a level where the reaction rate is negligible (disappearance). In this context, the total kinetic residence time is defined as the time from the time of introduction of the hydrocarbons into the apparatus to the quenching of the effluent, including the separation of solids from the reactants. The total conversion is therefore the sum of the catalytic reaction (time in contact with the catalyst) and the thermal reaction (time at reaction temperature). A typical FCC reaction environment has a relatively long residence time, including time for solids separation (usually greater than 1 second), and does not include quenching.
Field occurs at low temperatures under the above long residence times. Conversion is achieved at the lower temperatures due to extended contact with the catalyst. Thermal reactions are minimized at these low temperatures, thus reducing the need to quench the effluent. Although more C3 and higher compounds are produced compared to pyrolysis, the effluent has a disproportionately high concentration of paraffins due to increased hydrogen transport activity. Favorable conditions for olefin production, particularly high temperatures and short residence times, are particularly important when processing light feedstocks such as LPG and naphtha that require high temperatures to initiate and maintain the reaction catalyst or heat. difficult to achieve. All of the above methods improve the cracking of hydrocarbons into olefins. However, these processes either require large capital and operating costs due to fixed bed operations and hydrogen sulfide dilution, or reduce the yield of the desired olefin. Hydrogen sulfide used as a diluent also poses environmental and health problems due to its extremely high toxicity. Higher olefins, i.e. propylene, butene,
were found to be obtained in high yields by cracking hydrocarbons in a fluidized solids cracking environment with short residence times in the presence of acidic cracking catalysts alone or in combination with noble metal oxide dehydrogenation catalysts. It was.
This short residence time is achieved through a combination of a low residence time reactor, a very short residence time separation system, and quenching of the product. It is therefore an object of the present invention to provide a process by which hydrocarbons can be catalytically cracked to produce olefins and aromatic hydrocarbons. Another object of the present invention is to provide a method for preferentially cracking hydrocarbons to obtain C3-C5 olefins and/or aromatic hydrocarbons. Another object of the invention is to provide a method by which hydrocarbons can be cracked into desired products by varying the catalyst system. Another object of the present invention is to provide a reaction system that includes a quenching step to preferentially decompose hydrocarbons to obtain C3-C5 olefins and/or aromatic hydrocarbons while avoiding thermal degradation of the product. . SUMMARY OF THE INVENTION The present invention generally provides preferential decomposition of hydrocarbons to produce olefins, preferably C3-C5 olefins, and aromatic hydrocarbons at the acid sites of a catalyst solid, and optionally, the resulting paraffin isomers. The present invention relates to a method for catalytically dehydrogenating bodies to produce olefins. Acidic catalytic cracking of hydrocarbons proceeds by a carbonium ion mechanism, unlike the free radical mechanism of thermal decomposition. Carbonium ions are produced by extracting hydride ions from carbon-hydrogen bonds. Extraction of hydride ions and production of carbonium ions is catalyzed by acid sites on the catalyst solids. Carbonium ion decomposition also occurs at the beta position, thereby producing olefins and primary carbonium ions. The first carbonium ion undergoes a rapid ion shift (isomerization) to generate a second or third carbonium ion. This, combined with beta-decomposition regulation, produces high yields of propylene without simultaneously producing large amounts of ethylene. The ethylene found in the product is the result of competing free radical decomposition routes. In addition to providing a carbonium ion mechanism for isomerization, acid sites on the catalyst facilitate hydrogen transfer. At the temperatures contemplated by this invention, thermodynamic equilibrium favors olefins over paraffins, but increased hydrogen transfer activity results in a disproportionately large yield of paraffins. This is particularly true for branched isomers such as isobutylene. In these examples, if a particular dehydrogenation catalyst is used in combination with an acidic decomposition catalyst, the yield distribution shifts toward thermodynamic equilibrium and high concentrations of the desired olefin are obtained. In the present invention, kinetic residence time is defined as the total time from the point at which the hydrocarbons are introduced into the reaction zone to the point at which the cracked products are quenched, including the time for intermediate separation steps. This distinguishes the method of the invention from other methods in which the measurement of residence time is terminated before the point of separation and quenching. This is especially important. This is because catalytic cracking of hydrocarbons always proceeds in parallel with thermal cracking. The extent to which products are formed catalytically or thermally is a function of catalyst activity, catalyst usage, catalyst residence time, reaction temperature distribution, and total kinetic residence time in the thermo-catalytic environment. For example, mildly acidic catalyst activity at elevated temperatures can be used to shift the production of diolefins to paraffins and olefins without substantially changing the proportion of carbon products obtained by pyrolysis. Also, highly active acidic decomposition catalysts can be used at significantly lower temperatures to minimize thermal pathways and maximize acidic catalyst product distribution. Additionally, it has been discovered that catalytic dehydrogenation catalysts can be used in combination with acidic cracking catalysts to shift the reaction to suit olefin production. The present invention is particularly suitable for pyrolyzing hydrocarbon feedstocks such as C4 to C7 paraffins, naphtha, and light gas oils to higher olefins and/or aromatic hydrocarbons having 3 to 5 carbon atoms. It should be noted, however, that the process of the present invention is widely used to crack a wide range of hydrocarbons, from light distillates to heavy residual oils. The method of the present invention is to transfer preheated hydrocarbon feedstock and steam into a downflow tubular reactor.
Proceed by feeding the mixture to the top of the tube. At the same time, hot catalyst solid forms are introduced into the top of the reactor, and the hydrocarbons, steam, and catalyst solids combine and pass through the reactor zone, the separation zone, and the quench zone, during which the hydrocarbons undergo low severity and short residence times. It decomposes in time and the effluent is stabilized to prevent quality deterioration. Tubular reactor is about 900~1500〓, preferably
1000-1300〓 and operated at a pressure of about 10-100 pisa, the total motion residence time is about 0.05-2.0 seconds, preferably about 0.10-0.5 seconds. After being cracked from the cracked effluent, the catalyst solids are stripped of residual hydrocarbons, regenerated, and reheated in a transfer line and returned to the tubular reactor to continue the cracking process. The present invention is extremely well suited for use with short residence time fluidized solids crackers and short residence time separation devices, which devices are disclosed in U.S. Pat.
No. 4370303 (Woebcke et al.), U.S. Patent No. 4433984
(Gartside et al.), and U.S. Patent Application No. 084328
(Gartside et al.). The specific catalyst solids and catalyst to hydrocarbon ratio are selected based on the nature of the feedstock and the desired product distribution. Catalyst activity and catalyst loading will define the operating temperature at the short residence times employed in this invention, and thus determine the boundary between catalytic and pyrolysis reactions. The type of catalyst, acidic decomposition catalyst alone or in combination with noble metal oxidative dehydrogenation catalyst, will further determine the product distribution between olefins and paraffins. DETAILED DESCRIPTION OF THE INVENTION As mentioned above, the process of the invention is particularly suitable for producing olefins with high selectivity towards C3-C5 olefins and/or aromatic hydrocarbons.
The present invention is directed to a means of pyrolyzing hydrocarbon feedstocks in the presence of catalytically active heat-bearing solids. Hydrocarbons used as feedstocks include high boiling gas oils, atmospheric gas oils, naphtha, and C4 to C7 paraffins. However, the present invention enables catalytic cracking of a wide range of hydrocarbons to produce desired olefins and/or
or commonly used to produce aromatic hydrocarbons. Referring to the drawings, first, as shown in Figure 1,
The method of the present invention is carried out in a short residence time fluidized solids cracking system 1 (hereinafter referred to as QC system) comprising a tubular reactor 2, a reactor feeder 4, a separator 6, a quenching means 24, and a solids stripper 8. It will be carried out in The system 1 also includes means for regenerating the catalyst solids separated from the decomposition products after the reaction. The illustrated system is an entrained bed.
Heater 10, transport path 12, and fluidized bed vessel 14
including. In the heater, the catalyst solids are regenerated and reheated. In the fluidized bed vessel 14, the catalyst solids are removed from the combustion gases and then distributed again to the reactor 2. In operation, the hot catalyst solids are placed in a fluidized bed vessel 14.
from the reactor feeder 4 and mixes with the steam entering through line 16. The hydrocarbon feed enters preheater 20 through line 18 and then into line 2.
2 into the upper region of the tubular reactor 2.
Preheated hydrocarbon feedstock and catalyst solids and steam from reactor feed 4 travel through tubular reactor 2 . The hot catalyst solids, steam, and preheated hydrocarbons are in intimate contact within the reactor and cracking proceeds rapidly. After leaving the tubular reactor 2, the cracked hydrocarbon effluent and vapors are immediately separated from the catalyst solids in the separator 6, and the cracked effluent passes from the top to the quench zone 24, where it is The cracked products are immediately quenched by steam or light hydrocarbons which are transferred to quench zone 24 through quench line 26. This reduces the temperature of the mixture below the temperature at which a substantial thermal reaction occurs. The mixture of cracked products leaving the tubular reactor 2 and separated from the catalyst solids in the separator 6 may be quenched by passing through a bed of solids with catalytic (dehydrogenation) activity. Since the dehydrogenation reaction is endothermic, the mixture passing through will be cooled as the reaction progresses. In this case steam may be introduced to improve the reaction conditions. A preferred method of such quenching, as shown in FIG. Rapid cooling occurs here. The quenched product is sent to cyclone 28 where small amounts of catalyst solids are removed and sent through line 30 to solids stripper 8 where it is sent from separator 6 through line 32. Combines with most solids. In solids stripper 8, the catalyst solids are stripped of residual hydrocarbons by steam, nitrogen, or other inert gas fed to solids stripper 8 through line 34. The catalyst solids (which have carbon or coke deposits accumulated in the tubular reactor 2) are then sent to the heater 10. Air supplied to heater 10 through line 36 is mixed with the catalyst solids in heater 10 and this mixture is fed into transport channel 12 to transfer the catalyst solids to fluidized bed vessel 14. . The carbon precipitate on the catalyst solids is removed by combustion in the presence of air supplied from line 36 to provide the necessary heat for the cracking reaction. If additional fuel is required, it may be added into heater 10 from a fuel source (not shown). In essence, the method of the present invention involves a line 18 to a preheater 20 in which the temperature of the hydrocarbon is raised to about 800-900°C.
This is carried out by feeding a hydrocarbon, such as naphtha, atmospheric gas oil, or a mixture thereof, through the reactor. At the same time, catalyst solids from fluidized bed vessel 14 are sent to reactor feed 4 (best shown in FIG. 2) where catalyst solids are transferred to line 16.
It is mixed with the steam supplied through the reactor 2 and then sent to reactor 2 at a temperature of 1000 to 1600 °C. The feed ratio of catalyst solids to hydrocarbon is from 1 to 60:1 based on the weight of catalyst used. water vapor/
The feed ratio of hydrocarbons is 0:1.0, preferably 0.0:0.3
It is. Optionally, the catalytic cracking process may be initiated by injecting an alkane such as ethane through injection line 16 into tubular reactor 2 to form olefins and free radicals. This increases isomerization by forming carbonium ions and tends to stabilize the production of heavy hydrocarbons by competing with the generated free radicals. Such alkanes are added immediately upstream of the hydrocarbon feedstock 22. Catalyst solids suitable for this invention include silica gel,
It can be one of the commonly available supports with acidic properties such as alumina, clay. The catalytic solids can be associated with other catalytically active substances. Alternatively, the catalyst system used may be a conventional zeolite-based FCC catalyst, or one of the highly active ZSM-5, or a rare earth zeolite-based catalyst. The catalyst used can also include a dehydrogenation catalyst consisting of a noble metal oxide (such as an oxide of iron, chromium, platinum, etc.) supported on a suitable support such as silica alumina. Alternatively, the catalyst can be a mixture of the aforementioned catalysts in order to achieve a particular yield distribution. In tubular reactor 2, the mixed hydrocarbon feedstock is 800 to 1100 ml and the catalyst solid is
Heat to 1200 to 1700℃. The ratio of solids to hydrocarbons is set by the thermal balance and the desired catalyst activity. The effluent cracked products and catalyst solids effluent from tubular reactor 2 flows directly into separator 6 (detailed in FIG. 3). In this separator 6, separation into a gaseous product phase and a catalyst solids phase takes place. Gaseous products are line 2
4 and the catalyst solids are sent through line 32 to solids stripper 8. In-line quenching of the gaseous product is provided in quench zone 24 through quench line 26 . Cold solids, water,
Steam, light hydrocarbons, and reclaimed oil are examples of suitable quenching materials. Instead, the catalytic reactor 25
(see Figure 5), quenching is effected by passing the product through a catalyst bed and further reactions are carried out without the presence of solids. The total residence time from the time of introduction of the hydrocarbons into the tubular reactor 2 to the time of quenching in the quench zone 24 (optionally including the catalytic reactor 25) is preferably between 0.1 and 0.3 seconds. In the solids stripper 8, impurity gases are removed from the catalyst solids by a stream of steam, nitrogen or inert gas sent through line 34. Steam is removed from solids stripper 8 through line 30. The degassed catalyst solids are removed from the stripper 8 through line 38. The catalyst solids from the tubular reactor 2 with carbon deposits are transferred to an entrained bed heater.
Send to 10. In this heater 10, line 36
Air is passed through to provide the atmosphere necessary for regeneration of the catalyst solids. The catalyst solid is heater 1
0 and returned through transport line 12 to fluidized bed vessel 14. In the transport line 12, the catalyst solids are subsequently regenerated. In addition, regeneration of the catalyst solids can be carried out by transferring the catalyst to the fluidized bed vessel 1.
4, the temperature of the catalyst solids is raised to about 1200-1700°C. The details of reactant feeder 4 are fully described by Gartside et al., US Pat. No. 4,338,187;
It is incorporated herein by reference.
The reactant feeder of Gartside et al. has the ability to rapidly mix hydrocarbon feedstock and catalyst solids. As shown in FIG.
transfers the catalyst solids from a solids vessel or fluidized bed vessel 70 to the tubular reactor 2 through a vertically arranged conduit 72, and at the same time transfers the hydrocarbon feedstock to the tubular reactor 2.
It is transferred to the tubular reactor 2 by entering the path of the catalyst solids leaving the conduit 72 at an angle.
An annular chamber 7 fed with hydrocarbons by a single passage consisting of a toroidal feed line 76
4 terminates in an angled aperture 78. A mixing baffle or plug 80 also effectively aids in rapid and substantial mixing of the hydrocarbon feedstock and catalyst solids. The edge 79 of the angled opening 78 is connected to the edge 79 of the reactor end of the conduit 72.
It is preferable to have a convergent slope as shown in FIG.
In this manner, the gaseous hydrocarbon stream from chamber 74 is injected into the mixing region at an angle;
The catalyst solids phase flowing from conduit 72 is intercepted. The projection of the gas forms a cone indicated by dotted line 77, the vortex of which is below the solids flow path. By introducing the gaseous hydrocarbon phase at an angle, the two phases mix rapidly and uniformly to form a homogeneous reaction phase. Mixing of the solid and gas phases is determined by the shear surface and flow area between the solid and gas phases.
area). A case where the ratio of shear surface to flow area (S/A) is infinite is defined as complete mixing. By far the worst mixing occurs when solids are introduced into the walls of the reaction zone. In the system of this invention, the gas flow is introduced into the solid body in an annular manner, which ensures a large shear surface. As described in the preferred embodiment, phase infiltration is also obtained by adding a lateral gas flow through the annular feed means, resulting in even faster mixing. By using multiple annular gas supply points and multiple solids supply conduits,
Better mixing is promoted more quickly. This is because, for a constant solid flow area, the ratio of shear surface to flow area increases. Mixing is also a known function of the length to diameter ratio of the mixing region. The graph effectively reduces the diameter D while keeping the length L constant. For this reason,
Mixing increases. Plug 80 reduces the flow area and creates a separate mixing region. The combination of annular gas added around each solids feed point and a restricted separate mixing area greatly enhances the mixing conditions. Using this preferred embodiment, the time required to obtain an essentially homogeneous reaction phase within the reaction zone is very short. Therefore, this preferred gas and solids addition method can be used for reaction systems with residence times of less than 1 second, and even for reaction systems with residence times of less than 100 milliseconds. can. For the environment of the annular reactor 2 and reactant feeder 4, the walls are lined with an inner core 81 of ceramic material. As shown in FIG. 3, the separator 6 of the QC system can also be used to rapidly and separately separate the product and catalyst solids exiting the tubular reactor 2. The inlet to the separator 6 is directly above a right corner 90 that collects a large amount of catalyst solids 92 into a chamber 93. An optional weir 94 downstream of right angle corner 90 assists in depositing mass of solids 92, particularly when operating at a scale smaller than commercial scale production. The gas outlet 24 of the separator 6 is oriented 180° from the separator gas-solids inlet 96, and the solids outlet line 32 is directly opposite the gas outlet 24, with both the gas outlet line 24 and the weir 94 downstream of. During operation, centrifugal force is applied to the inlet 9 of the chamber 93.
The catalyst solids are allowed to advance to the wall facing the chamber 6, while the gas portion with low momentum is allowed to flow into the gas region of the chamber 93. Initially, the catalyst solids impinge on the wall facing the inlet 96 and then deposit to form a solid bed of solids 92. This solid bed forms a surface profile of approximately a 90° curved arc of a circle. The solids that hit the bed 92 move along a curved arc to the solids outlet 95. it is,
Preferably, it is directed by the downward flow of solids due to gravity. The exact shape of the arc depends on the particular separator shape and velocity, mass flow rate, bulk density,
Determined by inlet flow parameters such as particle size. Erosion is minimal because the force exerted on the incoming solids is directed against the solid bed 92 rather than on the separator 6 itself. Therefore, the separator efficiency, defined as solids removal from the gas phase pumped through outlet 97, is 150 ft./ft.
not adversely affected by high inflow velocities up to sec.
Separator 6 is operable over a wide range of dilute phase densities, preferably from 0.1 to 10 lbs./ft. The separator 6 of this invention can achieve an efficiency of about 90%, and in a preferred embodiment, the removal rate of catalyst solids is
More than 97% is obtained. For the height H given to the chamber 93,
It has been found that efficiency increases as the spacing of the 180° U-bends of inlet 96 and outlet 97 approaches the width of inlet 96. Thus, for a given height H, the efficiency of the separator 6 is increased by reducing the flow path and therefore the residence time. If the inner diameter of the outlet 96 is Di, then the distance CL between the center lines of the inlet 96 and the outlet 97 is:
Preferably, the distance between the center lines is not more than 4.0 (Di), and most preferably the distance between the center lines is between 1.5 and 2.5 (Di). Less than 1.5 (Di) allows for better separation, but fabrication difficulties often reduce the attractiveness of this embodiment. If the latter embodiment is desired, a single mold design may be necessary since the inlet 96 and outlet 97 are too close to be welded together. It has been found that the height H should be at least equal to 1.5 x Di or 4 inches high, whichever is greater. If H is less than Di or 4 inches, the incoming flow
, thereby tending to reintroduce solids into the gaseous product being pumped through outlet 97. The height H is preferably twice Di to obtain even greater separation efficiency. Although not particularly limited, if the height H is too large, the residence time will simply increase without any substantial increase in efficiency. The width W of the flow path shown in FIG.
Preferably it is 0.75 to 1.25 times Di, most preferably 0.9 to 1.10 (Di). Outlet 97 may have any size internal diameter (Dog). However, velocities greater than 75 ft./sec. result in erosion due to residual solids entering the gas. The inner diameter Dog of the outlet 97 is
The dimensions should be such that a pressure difference exists between the solid stripper 8 and the separator 6 shown in FIG. 1 so that a static height of the solids is created in the solids outlet passage 32. The static height of the solids in the solids outlet passage 32 forms a positive seal that prevents gas from entering the solids stripper 8. solid stripper 8
The magnitude of the pressure difference between the flow path 32 and the separator 6 is
The height of the solids therein is determined by the force required to move the solids in the bulk stream to the solids outlet 95. As the difference increases, the solid stripper 8
The net gas flow to is reduced. The solid body with gravitational momentum overcomes the difference, while the gas is preferentially pumped through the gas outlet 97. The inner diameter Dog of the gas outlet 97 is preferably the same as the inner diameter of the inlet 96 in order to make the velocity of the outlet less than or the same as the velocity of the inlet when one outlet is used. . FIG. 4 is a sectional view of the separator 6 taken along section 4-4 in FIG. The vertical side walls 101, 102 need to be straight or slightly arcuate as shown by dotted lines 101a and 102a. The flow path through the separator 6 is thus substantially rectangular in cross-section with a height H and a width W as shown in FIG. In the embodiment shown in FIG. 4, the shape of the flow path is defined by adjusting the coating width of the walls 101, 102. Alternatively, buffles, insert plates, weirs, or other means may be used. Similarly, walls 103 and 104 across the channel
may be similar in shape, but this is not essential. The separator shell and internal passages are preferably coated with an erosion-resistant coating 105, which is necessary if solids are impacted at high speeds. Commonly commercially available materials for corrosion-resistant coatings include Carborundum Precast Carbofrax D, Carborundum Precast Alfrax 201, or equivalents. When the separator 6 is used at high temperatures, a thermally insulating coating 106 is placed between the shell and the coating 105.
and the internal passageways and their respective erosion-resistant coatings. Details of separator 6 are fully described in US Pat. No. 4,288,235, the contents of which are included here. Additional aspects of the apparatus used in the practice of this invention are disclosed in the aforementioned US patent application Ser. No. 084,328, the contents of which are incorporated herein. Referring to the drawings, particularly FIGS. 6-9, there is shown a reactor system 204, a solids regenerator 20
8, and a system 202 that includes a solids evacuation system 210 is described. As shown in FIG. 7, reactor system 204
includes a central mixing section 211, an elongated reaction section 212, an extension 213 downstream of the elongated reaction section 212, a separator 206, and a quench system (shown in FIG. 8). The mixing section 211 is formed by a plug section 214 shown in cross section with an arcuate lower surface 215 . A horizontally placed plate 217 is placed above the plug part 214 with a space therebetween to form a solid inflow path 219 into the mixing part 211 . The solid inflow channel 219 has a cross section with right-angled corners, and has a rectangular opening 22.
5, where the solid particles enter the mixing section 21
1 in the form of a solid curtain 226. Horizontal inlet 22
5 is directly above each hydrocarbon feed port. A passage 203 having the shape of a ventilate extends from the solids inlet passage 219 to the hydrocarbon feed port 228 . A steam plenum (not shown) is disposed along the longitudinal edge of each horizontal inlet passageway 225 to direct preaccelerated gas (steam) from nozzles (not shown) through the horizontal openings 225 into the solid curtain 226. to be discharged. A gas discharge channel (not shown) is provided for discharging gas, usually steam or light hydrocarbons, under nozzle pressure. The nozzle is placed at a downward 45° angle to the horizontal. The preacceleration gas is discharged into the plenum at a pressure of 3 to 5 psi above the pressure within the reactor and is discharged through the nozzle at the same relative pressure at a rate of about 150 feet/second.
The preacceleration gas accelerates the flow of solids through the horizontal opening 225 from as little as 3 to 6 ft/sec to approximately 50 ft/sec due to mixing of the solids and preacceleration gas. A more detailed description can be found in U.S. Patent Application No. 084328
It has been in the issue. The hydrocarbon inlet 228 is located in the reactor wall located either perpendicular to the solid curtain 226 or at an angle of 30° above the solid curtain 226 . The hydrocarbon feedstock is transferred to manifold 22 via line 224.
It is released in 3. As can be seen in FIG. 7, the supply nozzles 228 are diametrically opposed to each other in the same horizontal plane. Reactor mixing zone 211
is rectangular and has a shape that is transferred to the tubular reactor in the elongated reaction section 212. The feedstock entering the mixing zone 211 through the nozzle 228 immediately impinges on the solids curtain, resulting in the desired mixing of the feedstock and hot particulate solids. Opposing sets of nozzles 228 cause opposing feedstock jets and entrained solids from solids curtain 226 to be directed by the precise contours of plug section 214 and impact each other at a generally vertical centerline of mixing zone 211. will give. When gas-solid mixed phase hydrocarbons are fed through the nozzle, the nozzle 228 connects the solids curtain 226 and 90
Arranged in degrees. When the hydrocarbon feedstock is a gas, the nozzle is oriented 30 degrees upward relative to the solids curtain. The majority of the amount of solids entering the mixing zone 211 of the reactor system 204 through the horizontal inlet 219 is in the mixing zone 211 of the reactor system 204 and in the solids control hopper 231 immediately above the horizontal inlet 219. Storage tank 21
8 and the upper chamber 231a. Pressure probes 233 and 235 are located in mixing zone 211 and control hopper chamber 231a, respectively, of reactor system 204 to measure pressure differences. The gas under pressure ( Steam) is line 230
Control hop perch yambar 231 through
released into a. As shown in FIG. 7, the separator 206 is
Mixed phase inlet 232, horizontal chamber section 2
34, with a plurality of cracked gas outlets 236 and particulate solids outlets 238. The basic principles regarding relative diameters (Di, Dog, Dos), chamber height (H) and length (L) described in the first embodiment are also applicable here. Separator 206 includes an elongated cracking zone 212 and reactor system 2
It is arranged in connection with the widening section 213 of 04. The flare section 213 terminates in a parator mixing phase inlet 232 centrally located at the top of the horizontal section 234. As a result of the configuration of the complex reaction system including separator 206, solid phase 242 is separated from bed 2 of horizontal section 234.
40, the cross-sectional profile 243 of phase 242 forms a curve along which the gases and solids of the mixed phase move. Expansion of the solids and cracked gas in expansion section 213 promotes heat transfer and velocity of the solid-gas mixture entering separator 206. The solids are routed to the side ends 246 of the horizontal section and are discharged downwardly through the solids outlet 238. The cracked gas passes through a 180 degree path and, after separation from the solids, passes through the side end 246 to the horizontal section 2.
is exhausted through a gas outlet 236 located at the top of 34. Multiple solids outlets 238 and gas outlets 236 simultaneously provide for both minimum time and maximum solid-gas separation in the separation zone. Separation or quenching system 207 also includes a conventional cyclone separator immediately downstream of each gas outlet 254, as shown in FIG.
The inlet line 254 to each cyclone separator 250 is oriented at a 90 degree angle to the gas outlet 236, and the cyclone separator 250 is oriented perpendicular to the system. Cyclone separator 250 serves to collect entrained solid particles discharged from separator 206. Diplets grain line 349 returns solid particles to regeneration assembly 208 and sends cracked gas downstream through gas outlet 251. Each cyclone inlet line 254 extending from cracked gas outlet 236 to cyclone 250 is provided with a direct quench line 252 . The quench oil, typically a 100-400 fraction, from the downstream distillation column is introduced directly into cyclone 250 through quench line 252 to terminate the reaction of the cracked gases. As shown in FIG.
is stripper 253, lift line 255,
and a recycled solids container 260. Stripper 253 is a tubular vessel into which solid particles from separator 206 are discharged through solids outlet legs extending from separator 238 and from cyclone dipleg 249. Ching 2 with nozzle opening 264
62 is provided at the bottom of the stripper 253. Stripping gas, typically steam, is discharged into ring 262 for exhaust through nozzle 264. Stripping steam rises through the layer of solid particles to remove impurities from the surface of the solid particles. The stripping steam and entrained impurities rise through the solid particles in the stripper 523 and are discharged through a vent line (not shown) to the cracked gas line. The stripped solids are transported to the entrained layer heater 2
Control hopper 2 for discharge to 58
55. control hopper 25
5 is a collection vessel into which the solids enter through standpipe 266 and from which an outlet line 273 extends to discharge the solids to entrained bed heater 258. The combination of control hopper 255 and standpipe 266 provides a falling bed solids transfer system. The pressure difference between the falling bed surface 268 in the control hopper and the outlet of the outlet line 273 determines the solids flow rate between the control hopper 255 and the entrained bed heater 258. A line 272 is provided to selectively introduce steam under pressure into the control hopper 255 to adjust the pressure differential. Probes 267 and 269 are respectively connected to control hopper 25 to monitor the pressure differential and adjust valve 265 in steam line 272.
5 and entrained layer heater 258. Entrainment layer heater 258 is essentially of tubular construction. Rows of independent fuel nozzles 261 supplied by fuel lines 263 are arranged approximately symmetrically on the lower sloped surface 275 of entrainment layer heater 258 . The pressurized air is passed through the entrainment layer heater 2 through a nozzle 277 arranged to cause the air to rise axially through the entrainment layer heater 258.
Enter 58. The air jet transports the solid particles through an entrainment bed heater 258 and into a recycled solids container 260.
It provides the driving force to raise the fuel, and the air necessary for combustion. The fuel ignites upon contact with solid matter in the presence of air. The combustion gas/solids mixture rising through lift line 257 enters regenerated solids container 260 tangentially, preferably at a 90 degree angle with the lift line, to separate the combustion gases from the solids. As shown in FIG. 6, a distube 285 is included within the gas outlet nozzle 286 to provide a cyclonic motion that improves the separation efficiency of the system. The recycled solids container 260 is connected to a standpipe 2 extending to the reactor hopper 231, as shown in FIG.
Equipped with 71. As shown in FIG. 9, the construction of the recycled solids container 260 provides for the accumulation of a falling layer 260, above which pressure is applied to the controlled release of recycled solids particles into the reactor hopper 231. is possible. The upper collection container 260 shown in FIGS. 6, 7, and 9 is
It includes a stripping section as a lower part with a stripping ring 279 to form the solids release system 210. Above the ring 271, the solids are fluidized and below the ring 271,
The solid material falls and the standpipe 271 shown in FIG.
is supplied to The standpipe 271 supplies a falling layer within the control hopper 231, as shown in FIG. The solids flow through standpipe 271 to reactor hopper 231, displacing the solids that were flowing to reactor 204. Once the entrance to the hopper 231 is covered, unaerated solids (fallen solids) will not continue to flow into the reactor hopper 231. in this way,
The location of inlet 231 defines the level of solids within hopper 231. As solids flow from the hopper due to the pressure difference between the space in the chamber 231a above the layer 218 and the mixing zone 211, the inlet 282 is uncovered and the hopper 2
Allow additional solids flow to 31. One embodiment of the invention shown in FIG. 1 is illustrated by the following comparative example (table). In this example, light FCC naphtha was processed by conventional tube pyrolysis to (0.9 seconds) by conventional catalytic cracking using a conventional active catalyst with an FCC residence time of greater than 1 second.
It was cracked by catalytic cracking with a high activity catalyst at FCC residence time and by very short residence time cracking plus quenching (QC system) with a similar high activity catalyst. Two cases with high activity catalysts are presented to show the effect of residence time on olefin yield.

TABLE In the table, Example A shows the yield obtained using conventional pyrolysis operated at typical pyrolysis temperatures and residence times. Example B shows a conventional catalytic start-up reactor using longer residence times and lower temperatures than the pyrolysis of Example A. As shown in the table, the conversion with the conventional catalyst is substantially lower than that used in the pyrolytic cracking of Example A. This low conversion is due to insufficient catalytic activity on this relatively light feedstock (816
516℃). However, even at such low conversion rates, the total C3
and C4 compounds are higher than for thermal decomposition as a result of the carbonium ion mechanism. Furthermore,
The ratio of C4 to C3 olefins is substantially increased due to the hydrogen transfer activity of the acidic cracking catalyst. Example C shows the yield of product that would be obtained by using a high activity acidic catalyst at low FCC residence time or high QC residence time without quenching. The selected operating conditions for Example C will result in reduced methane and ethylene yields compared to Example A. Conversion is increased compared to Example B even at lower temperatures (510°C versus 565°C) due to increased activity. There is a substantial increase in total C3 products as a result of the acidic decomposition catalyst, but the C4 yield decreases due to increased conversion. Additionally, due to the longer residence time, there is a substantial amount of hydrogen transfer as evidenced by the unacceptably high C3 paraffin/olefin compared to either Examples A or B. Example D shows a significant improvement in olefin yield by employing the method of the present invention in a very short residence time QC system. As shown in the table, when the reactor temperature increases by about 30℃ and the total kinetic residence time, i.e., decomposition reaction plus separation plus quenching, decreases to barely 0.15 seconds, the C3 olefin yield increases by about 100%. There is. In addition, the paraffin/olefin ratio is reduced to less than half of that obtained at the longer residence time of Example C. The paraffin/olefin ratio in this case is higher than in the case of pyrolysis at similar residence times as a result of the hydrogen transfer activity of the catalyst. However, the methane yield was lower than the lower level of Example C, and the C4 yield was improved by about 100%, indicating less secondary decomposition due to rapid cooling and short residence time reaction. . In other embodiments of the invention, the QC system comprises:
Used to facilitate the production of aromatics, especially benzene. The table shows the use of a QC system to accelerate the production of aromatics from the decomposition of n-hexane at a fixed conversion of 70%. Two examples of QC systems, one with a highly active catalyst and the other with a deactivated zeolite catalyst, are comparable to conventional pyrolysis.

[Table] It can be seen that in Example 1 of the table, a conversion rate of 70% of the hydrocarbon feedstock was obtained at a reactor outlet temperature of 730° C. and a residence time of 0.2 seconds. As shown in the table, pyrolysis produces substantial amounts of olefins, but only minor amounts of paraffins and aromatics.
Similar results are obtained when using completely inert solids such as pure alumina in QC cracking conditions. In a QC system using a highly active acidic cracking catalyst, a hydrocarbon conversion of 70% is obtained at a temperature as low as 550° C. and a residence time of 0.2 seconds (Example 2). Ethylene production can be suppressed, but
The yield of C3 olefin and paraffin is promoted.
Additionally, very small amounts of aromatics are produced. If a deactivated zeolite catalyst is used instead of a highly active catalyst, a completely different yield spectrum is obtained (Example 3). Deactivation of zeolite catalysts is usually the result of prolonged exposure to high temperatures and steam that disrupt the zeolite matrix. This results in a substantial reduction in catalyst surface area and therefore catalyst activity. Using zeolite catalyst
In the FCC unit, catalyst solids are removed and
Fresh catalyst is added to maintain activity. Such used solid matter is suitably used as a deactivated catalyst. The required reaction temperature is similar to that required to achieve thermal decomposition conversion due to the low activity. However, quite unexpectedly, there is a substantial increase in aromatic (particularly benzene) production and a decrease in C3 and C4 olefin production. Yields of ethylene are similar to those obtained by pyrolysis dominated by free radical cracking reactions at these temperatures.
However, deactivated catalysts lead to accelerated aromatization activity at these higher temperatures and aromatics are formed under consumption of C3 and C4 olefins and paraffins. These unexpected results suggest that depending on the desired product, available feedstock, and catalyst selection, one may choose between high C3-C5 olefin production or high aromatics production according to the present invention. Allows the operator to change the operating conditions of the QC system. In other embodiments of the invention, a dehydrogenation catalyst is combined with an acidic decomposition catalyst. The examples shown in the table below use minas naphtha feedstock and compare catalytic cracking and thermal cracking. Catalytic cracking requires lower temperatures to achieve a given conversion;
Only 75% of thermal decomposition products (C1 and C2 compounds)
%. Carbonium ion decomposition will shift the yield spectrum in favor of C3 and C4 compounds. As shown in the table below, the use of only acidic catalysts reduces the total isoC4 due to the ionic nature of the decomposition.
resulting in a substantial increase in (paraffin + olefin). However, many of the isocompounds are isoparaffins, and thermodynamic equilibrium favors the formation of olefins rather than paraffins. In the case of coil pyrolysis, iso and normal C4 are 87% olefinic, indicating an approach to equilibrium. However, the normal compound is 70% olefinic and the iso compound is 29% olefinic.

【table】

[Table] This is because the hydrogen transfer activity of the catalyst leads to a new equilibrium relationship based on reaction kinetics rather than thermodynamics. Iso compounds show a greater tendency to undergo hydrogen transfer in the presence of acidic catalysts than normal compounds. When a noble metal oxide catalyst is added to the flowing acidic decomposition catalyst, the paraffins produced in the acidic sites are dehydrogenated to olefins. The extent to which this reaction occurs depends on the concentration and activity of the dehydrogenation catalyst. In the table, tin and chromium oxides are mixed with acidic decomposition catalysts to achieve 80% approach to equilibrium for iso compounds and 87% approach to equilibrium for normal compounds.
As shown in the table, the production of normal and iso valuable C4 olefins is substantially increased. The production of isobutylene from the same raw material increases by a factor of 3 or more, and the production of normal butylene increases by a factor of 2 or more. The use of mixed catalyst systems provides additional product distribution flexibility to the catalytic process. Rather than mixing the dehydrogenation catalyst and the acidic decomposition catalyst and causing a mixed reaction by quenching, for example using steam, a packed bed in the contact reactor 25 located downstream of the main separation in the separator 6 is used. A dehydrogenation catalyst can be located. Paraffins formed upon contact with acidic catalysts will be dehydrogenated to olefins.

[Brief explanation of the drawing]

Fig. 1 is a schematic diagram of the process of the present invention, Fig. 2 is a longitudinal cross-sectional view of a reactor feeder used in the apparatus of the present invention, Fig. 3 is a longitudinal cross-sectional view of a separator employed in the present invention, and Fig. 4 is a 4-4 sectional view in Figure 3,
FIG. 5 is a schematic diagram of the quenching process of the present invention, FIG. 6 is a front view of the entire system according to one embodiment of the present invention, and FIG. 7 is a longitudinal cross-section of the reactor and gas-solid separator employed in the present invention. FIG. 8 is a cross-sectional view taken along line 8-8 of FIG. 7, and FIG. 9 is a front view of the solids recycling assembly employed in the present invention. 1... cracking system, 2... reactor,
4... Reactor feeder, 6... Separator,
8... Stripper, 10... Heater, 12...
...Transportation line, 14...Fluidized bed container, 20...Preheater.

Claims (1)

[Scope of Claims] 1. A method for producing an offrein, an aromatic compound, or a combination thereof by catalytically cracking a hydrocarbon feedstock, comprising: (a) introducing the hydrocarbon feedstock into a cracking reactor; (b) simultaneously transferring a thermally acidic decomposition catalyst solid to the reactor; (c) catalytically cracking the hydrocarbon feedstock using heat supplied from the thermally acidic catalyst solid to produce a decomposition product; and (d) decomposing the hydrocarbon feedstock. the residence time of the hydrocarbon feedstock from step (a) to step (e), including the steps of separating products from the thermal catalyst solids and (e) quenching the separated cracked products; is within a range of about 0.1 to 2.0 seconds. 2. The method of claim 1, wherein the residence time is about 0.05 seconds to 0.5 seconds. 3(f) transferring said separated catalyst solids to a stripper to remove residual cracked gas products; and (g) combusting said separated catalyst solids, thereby removing carbon deposits. 2. The method of claim 1, further comprising the steps of: heating the stripped catalyst solids to produce regenerated catalyst solids; and (h) transferring the regenerated catalyst solids to the cracking reactor. . 4. The method of claim 1, wherein a portion of the catalyst solids comprises a dehydrogenation catalyst. 5. The method of claim 4, wherein the dehydrogenation catalyst is selected from noble metal oxides on an inert support. 6. The method of claim 1, wherein the temperature of the catalytic decomposition reaction is about 900-1500°C, and the weight ratio of catalyst solids to hydrocarbon feedstock is 1-60°. 7. The temperature of the catalytic decomposition reaction is about 1000~1300
The method according to claim 6, wherein the residence time is 0.1 to 0.3 seconds. 8 the olefin has 3 to 5 carbon atoms,
2. The method of claim 1, wherein the aromatic compound has 6 to 8 carbon atoms. 9. The method of claim 1, wherein the hydrocarbon feedstock is selected from C4-C7 paraffin, naphtha, and light gas oil. 10. The method of claim 1, wherein the active acidic catalyst solid is a zeolitic catalyst. 11. The method of claim 10, wherein the inert carrier is selected from the group consisting of silica gel, silica alumina, viscosity, and mixtures of any of these. 12. The method of claim 11, wherein the decomposition products are primarily monoaromatic compounds and the acidic catalyst solid is thermally deactivated. 13 transferring the hydrocarbon feed stream and thermal catalyst solids to a tubular thermal regeneration cracking reactor through a feeder, the feeder having a longitudinal passage communicating with the cracking reactor and the thermal solids in a thermal solids vessel; to locally fluidize the thermal catalyst solids above the longitudinal passageway and transport the hydrocarbon feedstock into the thermal regeneration reactor at an angle to the path of the catalyst solids entering the thermal regeneration reactor. 2. The method of claim 1, including the step of providing. 14 separating the thermal catalyst solids and cracked product gas in a separator, wherein the catalyst solids and cracked product gas enter the separator through a separator inlet and after changing direction by 90 degrees; , the product gas is further redirected by 90 degrees, such that the product gas is redirected by 180 degrees from the inlet, while the catalyst solids remain in a passage redirected by 90 degrees from the inlet; A method according to claim 1, characterized in that the passage of the catalyst solids is then directed downwards and the separated product gas is quenched. 15 separating the catalyst solids and decomposed gas from within a separator, wherein the separator separates the solids from the mixture of solids and gas by about 80%;
% rapid separation, the chamber having vertical side walls surrounded by straight lines, the side walls being approximately rectangular in cross-section and forming a flow channel of height H and width W, and the chamber or the It has an inlet of internal diameter Di for the mixture, a gas outlet, and a solids outlet, said inlet being arranged at one end of said chamber perpendicular to said flow channel of said height H, said height H being at least said internal diameter Di. 4 inches greater than or equal to
0.75Di to 1.25Di, said solids outlets are located at opposite ends of said chamber and are arranged to force solids to drain downwardly by gravity, and said gas outlets are arranged to located at a distance of less than 4D from the inlet and arranged to change the direction of gas flow by 180 degrees, the resulting centrifugal force directing the flow of said incoming catalyst solids in the direction of the side wall opposite said inlet. , and maintaining a substantially stationary bed of solids, the surface of said stationary bed forming a curvilinear path of about a 90 degree arc of a circle to flow the solids to said solids outlet. The method according to claim 1.