US20030116471A1

US20030116471A1 - Catalytic cracking process of petroleum hydrocarbons

Info

Publication number: US20030116471A1
Application number: US10/229,155
Authority: US
Inventors: Ruichi Zhang; Xiaoxiang Zhong; Jiushun Zhang; Kejia Xu; Zhigang Zhang; Shuandi Hou; Xueliang Chang; Zhiguo Wu
Original assignee: China Petroleum and Chemical Corp
Current assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Priority date: 2001-08-29
Filing date: 2002-08-28
Publication date: 2003-06-26
Also published as: FR2829143A1; FR2829143B1; JP4412886B2; JP2003105346A

Abstract

The present invention discloses a catalytic cracking process of petroleum hydrocarbons and a tube-in-tube riser reactor used therein. In the present invention, a catalyst is fed via an inlet conduit to an inner tube and an annular space between the inner and outer tubes of the tube-in-tube riser reactor to contact a hydrocarbon feedstock, and the hydrocarbon feedstock is cracked under FCC conditions, then a reaction stream thus produced flows into a separation apparatus via a confluence tube to separate a hydrocarbon product stream from a spent catalyst, the spent catalyst is stripped and regenerated, then a regenerated catalyst is recycled for reuse. The process of the present invention can improve both product distribution and product properties.

Description

TECHNIQUE FIELD

The present invention relates to a catalytic cracking process of petroleum hydrocarbons ire the absence of hydrogen

TECHNIQUE BACKGROUND

Although catalytic cracking process has been developed for several decades and formed a relatively complete system of techniques, refiners are still studying and searching untiringly to seek for a process not only meeting the requirements of increasingly rigorous environmental laws and regulations but also according with the change of market demands so as to create satisfactory economic benefit for enterprises.

U.S. Pat. No. 5,043,522 and U.S. Pat. No. 5,846,403 relate to the improvements of conventional catalytic cracking process, in which part of catalytic gasoline is fed to a riser reactor at upstream of the feedstock nozzles to contact a regenerated catalyst having high activity at a high temperature, so that the yields of light olefins such as propylene, butylene and the like are increased, meanwhile the octane number of gasoline is improved.

CN 1279270 A discloses a process for increasing simultaneously yields of both liquefied petroleum gas and diesel oil. In this process, a catalytic gasoline is also fed to a riser reactor at upstream of the feedstock nozzles to contact at first and then react with a regenerated catalyst. This part of reprocessed catalytic gasoline is cracked fully at a high temperature under a large catalyst-oil ratio to form a great quantity of liquefied petroleum gas. Meanwhile, coke is deposited on the catalyst in a minute quantity, so that the catalyst activity is reduced appropriately in favor of producing more diesel oil,

U.S. Pat. No. 3,994,933 discloses a catalytic cracking process with two riser reactors sharing a disengager In this process, a light cycle oil is fed to a riser reactor to contact a regenerated catalyst, resulting in a conversion less than 30%; the spent catalyst is fed to another riser to contact a fresh feedstock and heavy cycle oil.

CN 1069054 A discloses a flexible and multi-purpose catalytic cracking process of hydrocarbons. This process relates to two independent riser reactors and two sedimentation separation systems attached with the risers. In the first riser reactor, a light hydrocarbon is contacted and reacted with a regenerated catalyst under conditions of a temperature of 600-700° C., a catalyst-oil ratio of 10-40, a reaction time of 2-20 seconds, and the carbon content of the catalyst being controlled in the range of 0.1-0.4% by weight; the spent catalyst is fed to the other riser to contact heavy hydrocarbons under conventional catalytic cracking reaction conditions.

Both U.S. Pat. No. 4,820,493 and CN 1174094 A expand the diameter of the pre-lifting section of the riser reactor, in which an inner tube for transferring catalyst is set coaxially to improve the contacting effect of oil and catalyst and to increase yield of desired product. Said inner tube for transferring catalyst is located at a place below the feed nozzle of hydrocarbon oil Both U.S. Pat. No. 4,310,489 a CN 1096047 A use a catalytic cracking unit with two risers. Therein, a riser is used to treat a heavy feedstock of conventional catalytic cracking, the other riser is used for catalytic modifying poor diesel oils or heavy gasoline fractions.

In sum, the catalytic conversion processes for treating simultaneously light oil and heavy oil disclosed in the prior art, can be essentially divided into two categories: (1) using a single riser reactor and charging the light feedstock to the upstream point of heavy feedstock inlet, and (2) using two riser reactors and treating different feedstocks in different risers. The first category of the processes need a little modification on equipment, however, reaction conditions of light oils are basically fixed, and product distribution and product properties can hardly be improved by means of optimization of operation variables. The second category of the processes overcomes disadvantages of the first category of processes, operation conditions of each riser can be adjusted independently to carry out each reaction under respective conditions adapted to different feedstocks. However, for the second category of the processes, costs in both construction of unit and reconstruction of equipment are increased significantly. Furthermore, in practical industrial production, a complicated flowchart will greatly increase the difficulty in operation.

SUMMARY OF THE INVENTION

An object of the present invention is to provide a catalytic cracking process of petroleum hydrocarbons using a tube-in-tube riser reactor, which can provide hydrocarbon feedstocks having different properties with suitable reaction conditions and improve obviously the product distribution and product properties of the catalytic cracking process.

Another object of the present invention is to provide a relay cracking process of petroleum hydrocarbons, which can make the catalytic cracking process not only have a desirable ability to convert heavy oils and better product selectivity but also simplify flowchart and easy to operate.

The catalytic cracking process provided in the present invention comprises the following steps

(1) feeding a catalyst from an inlet conduit to an inner tube and an annular space between inner and outer tubes of a tube-in-tube riser reactor, which flows upward under an action of pre-lifting media

(2) feeding a hydrocarbon feedstock into the inner tube and the annular space between the inner and outer tubes of the reactor, which contacts the catalyst therein to form an oil-catalyst mixture, so that the reaction of the hydrocarbon feedstock is carried out under catalytic cracking reaction conditions to form a reaction stream which flows upward along vessel wall;

(3) the reaction stream both from the inner tube and from the annular space between the inner and outer tubes flows together at the inlet of a confluence tube, and then enters a separation apparatus via the confluence tube, where a hydrocarbon product stream is separated from a spent catalyst;

(4) further separating the hydrocarbon product stream into various products including gasoline, diesel oil and liquefied petroleum gas, stripping and regenerating the spent catalyst, and recycling a regenerated catalyst into the reactor for reuse.

Another catalytic cracking process of petroleum hydrocarbons provided in the present invention (namely relay cracking process) mainly comprises the following steps:

(1) sending a regenerated catalyst to the bottom of the tube-in-tube riser reactor via a catalyst conduit, which flows upwards under an action of a pre-lifting media, 20-80% by weight of the regenerated catalyst flowing into the inner tube, and the remaining part of the regenerated catalyst entering the annular space between the inner and outer tubes, which flows upward under an action of the pre-lifting media;

(2) feeding a hydrocarbon feedstock to the inner tube of the reactor to contact the catalyst therein and form a oil-catalyst mixture, so that the reaction of the hydrocarbon feedstock is carried out under catalytic cracking reaction conditions to form a reaction stream which flows upward along vessel wall;

(3) the reaction stream from the inner tube and the regenerated catalyst (stream) from the annular space flow together in the confluence tube, and making the reaction stream react continuously under catalytic cracking conditions; introducing the resulting reaction stream to a separation apparatus via the confluence tube, where a hydrocarbon product stream is separated from a spent catalyst;

(4) separating the hydrocarbon product stream further into various products including gasoline diesel oil and liquefied petroleum gas; stripping and regenerating the spent catalyst, and recycling the regenerated catalyst into the reactor for reuse.

In comparison with the prior art, beneficial effects of the present invention are mainly embodied in the following aspects:

The processes provided in the present invention are simple in equipment and flexible in operation. Not only the reactions of heavy oils and light oils can be carried out separately in their respective reaction zones, but also reaction conditions can be regulated flexibly according to physical-chemical properties and mass flow of different feedstocks, which thus create favorable conditions for improving product distribution and product quality.

The processes provided in the present invention can make flexible arrangements of several production schemes such as, for example, gasoline scheme, diesel oil scheme, liquefied petroleum gas scheme, light olefins scheme and the like. Therefore, petroleum refineries can adjust product distribution pattern in time by using the processes of the present invention in accordance with variation of market demand so as to obtain more profitably economic benefit.

The processes of the present invention can obviously improve catalytic cracking product distribution, reduce yields of dry Gas and coke, increase yields of high value products such as liquefied petroleum as, gasoline and/or diesel oil and the like.

Furthermore, the processes provided in the present invention can also improve the product quality and reduce the environmental pollution caused by petroleum products It is proved by test that the processes can decrease olefin content of gasoline and increase the octane number of gasoline, reduce freezing point of diesel oil., improve sensibility of the diesel oil to flow improvers, increase stability of the diesel oil; meanwhile, the processes have a certain effect in reducing contents of the impurities such as sulfur, nitrogen and the like in gasoline and diesel oils.

ILLUSTRATION OF FIGURES

FIG. 1 is a schematic structure diagram of a tube-in-tube riser reactor having a single conduit for feeding catalyst. [0026]
FIG. 2 is a schematic structure diagram of joint regions for an inner tube, an outer tube and a confluence tube in a tube-in-tube riser reactor having a single conduit for feeding catalyst. [0027]
FIG. 3 is a schematic diagram of setting modes of feed nozzles of an inner tube and an outer tube in a tube-in-tube riser reactor having a single conduit for feeding catalyst. [0028]
FIG. 4 is a schematic structure diagram of a tube-in-tube riser reactor having two conduits for feeding catalyst. [0029]
FIG. 5 is a schematic structure diagram of joint regions for an inner tube, an outer tube and a confluence tube in a tube-in-tube riser reactor having two conduits for feeding catalyst. [0030]
FIGS. [0031] 6-9 are principle flowcharts of the catalytic cracking processes of petroleum hydrocarbons using a tube-in-tube riser reactor
FIGS. 10 and 11 are principle flowcharts of the relay cracking process of petroleum hydrocarbons.[0032]

DETAILED DESCRIPTION OF THE INVENTION

The tube-in-tube riser reactor is used both in the catalytic cracking process of petroleum hydrocarbons and in the catalytic relay cracking process of petroleum hydrocarbons of the present invention. Said tube-in-tube riser reactor may have a single conduit or two conduits for feeding catalyst, or other reactors having similar structure. The structure of the tube-in-tube riser reactor is illustrated in detail hereinbelow in combination with the Figures [0033]
In the present invention, said tube-in-tube riser reactor with a single conduit for feeding, catalyst has a structure shown in FIG. 1. The reactor mainly includes the following members, regenerated [0034] catalyst inlet conduit 1, inner tube 2, outer tube 3, confluence tube 4, pre-lifting distribution rings 5, 6 and 7, and feed nozzles 8 and 9; wherein the inner tube 2 and outer tube 3 are coaxial, and the ratio of the cross-section area of the inner tube to the cross-section area of the annular space between the inner and outer tubes is 1:0.1-10, preferably 1:0.2-2, the lower end of the inner tube 2 is located at a place above the inlet of regenerated catalyst, the inner tube has a length amounting to 10-70% of the total length of the reactor, preferably 20-60%; one end of the confluence tube 4 is connected with the upper end of the outer tube 3, and the other end is connected with a gas/solid separation apparatus, the cross-section area ratio of the confluence tube 4 and the inner tube 2 is 1:0.2-0.8; the pre-lifting distribution rings 5, 6 and 7 are located at the bottoms of the reactor, inner tube and outer tube respectively. Fixing member 10 may be installed in multi-lines, for example, with 2-12 lines of shrouding wires or draw-bars between the inner and outer tubes according to the specific size of the reactor and requirement in engineering.
In the tube-in-tube riser reactor having a single conduit for feeding catalyst, the distance from the upper end of the inner tube (i.e. the outlet end of the inner tube) to the outlet end of the confluence tube is 1-30 meters (i.e., total length of the confluence tube), preferably 2-20 meters. Different types of the joint modes between the inner tube, the outer tube and the confluence tube may be set in accordance with requirements. FIG. 2 exemplifies four embodiments of the joint modes in the present invention, but does not intend to limit the present invention. [0035]
As shown in FIG. 2, in the embodiment A, the upper end of the inner tube is in a straight tube shape, and the outer tube is also in a straight tube shape; both the inner diameters of the confluence tube and the outer tube are the same. The inner tube has an inner diameter of D[0036] 1, the confluence tube has an inner diameter of D2, and D1:D2=0.4-0.9:1.
As shown in FIG. 2, in the embodiment B, the upper outlet section of the inner tube is diverged in diameter, the outer tube is a straight tube, and both the confluence tube and the outer tube are equal in inner diameters The inner tube has an inner diameter of D[0037] 1 and the confluence tube has an inner diameter of D2. The ratio of the height H1 of the upper outlet section of the inner tube to the inner diameter D1 of the inner tube is 0.5-3:1 The divergence angle a of the upper outlet section of the inner tube is 5-3°.
As shown in FIG. 2, in the embodiment C, the upper outlet section of the inner tube is diverged in diameter, and the outer tube is converged in diameter and then connected with the confluence tube. The inner tube has an inner diameter of D[0038] 1 and the confluence tube has an inner diameter of D2, The ratio of the height H2 of the upper outlet section of the inner tube to the inner diameter D1 of the inner tube is 0.5-3:1. The divergence angle b of the upper outlet section of the inner rube is 5-30°. The bottom end of the converged section of the outer tube is located between 1.0 D2 above the inner tube outlet and 1.0 D2 below the inner tube outlet with the convergence angle c of 10-60°. The ratio of the height H3 of the converged section of the outer tube to the inner diameter D2 of the confluence tube is 0.5-3.0:1.
As shown in FIG. 2, in embodiment D, the upper end of the inner tube is in a straight tube shape, and the outer tube is converged in diameter and then is connected with the confluence tube. The inner tube has an inner diameter of D[0039] 1 and the confluence tube has an inner diameter of D2. The bottom end of the converged section of the outer tube is located 10.5 between 1.0 D2 above the inner tube outlet and 1.0 D2 below the inner tube outlet with the convergence angle d of 5-45°. The ratio of the height H4 of the converged section in the outer tube to the inner diameter D2 of the confluence tube is 0.5-3.0:1.
Three kinds of designs for diverged, equal and converged diameters may be used to the lower end of the inner tube in the tube-in-tube riser reactor having a single conduit for feeding catalyst For example, in FIG. 3, the ratio of the height H[0040] 5 of the inlet section at lower part of the inner tube to the inner diameter D1 of the inner tube is 0.5-3:1 with the divergence angle e of −30-30°. Three kinds of designs for diverged, equal and converged diameters may be also used to the lower end of the outer tube, in which the specific design concept is similar to the design of the pre-lifting section in a conventional catalytic cracking riser reactor
As shown in FIG. 1, feed [0041] nozzle 8 of the tube-in-tube riser reactor having a single conduit for feeding catalyst may be installed at lower part of the inner tube at a place 5-30% of its total length, also, 2-4 layers of feed nozzles may be set along the inner tube. Or, as shown in FIG. 3, the feed nozzle may be set along central axis of the reactor so that it passes through the lower part of the outer tube and then gets into 0.1-3.0 meters in the inner tube. Feed nozzle 9 may be set at the lower part of the out tube at a place 5-30% of its total length; also, 2-4 layers of feed nozzles may be set in vertical direction of the outer tube. Feed nozzles 8 and 9 may be used in any form favorable to disperse homogeneously hydrocarbon feedstock. Feed nozzles 8 or 9 may consist of 2-12 feed nozzles arranged uniformly along circumferential direction. A quenching media nozzle may be set at the confluence tube in order to create conditions for injecting a quenching media. In a word, a relatively flexible feeding mode may be used, and multi-point feed injection may be performed in the inner tube and outer tube.
The tube-in-tube riser reactor having two conduits for feeding catalyst has a structure shown in FIG. 4. The reactor mainly includes the following members: [0042] catalyst inlet conduits 21 and 22, inner tube 35, outer tube 36, confluence tube 38, pre-lifting distribution rings 31 and 33, arid feed nozzles 32 and 34. Therein, inner tube 35 and outer tube 36 are coaxial, and the ratio of cross-section area of the inner tube to cross-section area of the annular space between the inner and outer tubes is 1:0.1-10, preferably 1:0.1-2 The catalyst inlet conduit 21 is connected with the lower end of the inner tube 35, which has a length of 10-70% of total length of the reactor, preferably 15-50%. A distance from the lower end of the outer tube 36 to the lower end of the inner tube 35 is 2-20% of total length of the reactor, preferably 5-15%. The catalyst inlet conduit 22 is connected with the lower end of the outer tube 36. One end of the confluence tube 38 is connected with the upper end of the outer tube 36, and the other end is connected with a gas/solid separation apparatus. A cross-section area ratio of the confluence tube 38 to the inner tube 35 is 1:0.2-0.8, preferably 1:0.3-0.7. The pre-lifting distribution rings 31 and 33 are located at the bottom of the inner tube and the bottom of the annular space between the inner and outer tubes respectively. Feed nozzles 32 and 34 are located at the lower parts of the inner tube and the outer tube respectively. Fixing member 30 may be installed in multi-lines, for example, with 2-12 lines of shrouding wires or draw-bars between the inner and outer tubes according to specific size of the reactor and requirements of the engineering.
In the tube-in-tube riser reactor having two conduits for feeding catalyst, the distance from the upper end of the inner tube (i.e. the outlet end of the inner tube) to the outlet end of the confluence tube (i.e. total length of the confluence tube) is 0 5-20 meters, preferably 1-10 meters. It is desirable to optimize the inner diameter of the confluence tube according to the linear velocity of oil-gas in the confluence tube, and the linear velocity of the oil-gas in the confluence tube should be 0.5-2.0 times of the linear velocity at the inner tube outlet, preferably 0.8-1.2 times. The initial apparent linear velocity of the pre-lifting media is 0.3-8 m/s in the inner tube, and the apparent linear velocity of the pre-lifting media is 0.2-10 m/s in the outer tube. [0043]
Connection regions of the inner tube, the outer tube and the confluence tube may be used with different ways according to requirements. FIG. 5 shows four modes of embodiments, which however do not intend to limit the modes used for the reactor. [0044]
As shown in FIG. 5, in embodiment A[0045] 1, the upper end of the inner tube is in a straight tube shape, the outer tube is also in a straight tube shape, and the confluence tube and outer tube are equal in inner diameter. The inner tube has an inner diameter of M1, and the confluence tube has an inner diameter of M2, and M2: M1=1.5-5:1.
As shown in FIG. 5, in embodiment B[0046] 1, the upper outlet sects of the inner tube is diverted in diameter, the outer tube is a straight tube, and the confluence tube and outer tube are equal in inner diameter. The inner tube has an inner diameter of M1 and the confluence tube has an inner diameter of M2 The ratio of the height N1 of the upper outlet section of the inner tube to the inner diameter M1 of the inner tube is 0.5-3. 1. The divergence angle a1 of the upper outlet section of the inner tube is 5-30°.
As shown in FIG. 5, in embodiment C[0047] 1, the upper outlet section of the inner tube is diverged in diameter, and the outer tube is connected with the confluence tube after converged in diameter. The inner tube has an inner diameter of M1 and the confluence tube has an inner diameter of M2. The ratio of the height N2 of the upper outlet section of the inner tube to the inner diameter M1 of the inner tube is 0.5-3:1. The divergence angle b1 of the upper outlet section of the inner tube is 5-30°. The bottom end of the converged section in the outer tube is located between 1.0 M2 above the inner tube outlet and 1.0 M2 below the inner tube outlet with the convergence angle c1 of 10-60°. The ratio of the height N3 of the converged section of the outer tube to the inner diameter M2 of the confluence tube is 0.5-3.0:1.
As shown in FIG. 5, in embodiment D[0048] 1 the upper end of the inner tube is in a straight tube shape, and the outer tube is connected with the confluence tube after converged in diameter. The inner tube has an inner diameter as M1 and the confluence tube has an inner diameter as M2 The bottom end of the converged section of the outer tube is located between 1.0 M2 above the inner tube outlet and 1.0 M2 below the inner tube outlet with the convergence angle d1 of 5-45°. The ratio of the height N4 of the converged section, in the outer tube to the inner diameter M2 of the confluence tube is 0.5-3.0:1.
The design of both lower ends of the inner tube and the outer tube in the tube-in-tube riser reactor having two conduits for feeding catalyst is basically similar to that of the pre-lifting section in a conventional catalytic cracking riser reactor. The design of the pre-lifting distribution rings [0049] 31 and 33 are also basically similar to that in a conventional catalytic cracking process.
In the tube-in-tube riser reactor having two conduits for feeding catalyst, the [0050] feed nozzle 32 is set at a position of 5-30% of total inner tube length along the lower part of the inner tube Feed nozzle 34 is set at a position of 5-30% of the total outer tube length along the lower part of the outer tube. Feed nozzles 32 and 34 may be used in any form, provided that it is favorable to disperse hydrocarbon feedstock homogeneously. Both feed nozzles 32 and 34 may consist of 2-12 feed nozzles, and said nozzles should be arranged uniformly along circumferential direction.
In the processes provided in tile present invention, the catalysts fed via catalyst inlet conduit respectively to the inner tube and the annular space between the inner and outer tubes of the tube-in-tube riser reactor may be the same or different. Specifically speaking, the catalysts fed to the inner tube and the annular space between the inner and outer tubes of the tube-in-tube riser reactor may be the regenerated catalysts from a regenerator at high temperature. Also, the catalysts may be a mixture of the regenerated catalyst and a spent catalyst and/or a semi-regenerated catalyst Or else, the regenerated catalyst may be fed to the inner tube of the tube-in-tube riser reactor, and the semi-regenerated catalyst or spent catalyst or a mixture thereof may be fed to the annular space between the inner tube and outer tube, and vice versa. In a word, the catalyst fed to the catalyst inlet conduit may be contemplated comprehensively and rearranged flexibly in accordance with the unit conditions, feedstock properties and requirement for the desired products. Furthermore, the catalysts to be fed to the tube-in-tube riser reactor may also be cooled via a catalyst cooler, or may be fed to the tube-in-tube riser reactor via catalyst inlet conduit after the regenerated catalyst and spent catalyst and/or semi-regenerated catalyst are fully mixed in a catalyst blending vessel. [0051]
The catalysts used in the present invention may be any catalyst suitable for catalytic cracking process, of which active components are at least one kind of zeolite selected from a group consisting of Y-type or HY-type zeolite containing or not containing rare earth and/or phosphorous, ultra-stable Y-type zeolite containing or not containing rare earth and/or phosphorous, ZSM-5 family zeolite or high-silica zeolite having a pentasil structure, β-zeolite, ferrierite, and also an amorphous alumina-silica catalyst. The inner tube and the annular space between the inner and outer tubes may use two different kinds of catalysts. [0052]
In the processes provided in the present invention, the bottom of the tube-in-tube riser reactor, the bottom of the inner tube and the bottom of the annular space between the inner and outer tubes are furnished with an inlet of a pre-lifting media respectively. Steam, dry gas or a mixture thereof can be used as a pre-lifting media. The apparent linear velocity of the pre-lifting gas in the inner tube is initially 0.3-6.0 m/s, and the apparent linear velocity of the pre-lifting gas in the annular space between the inner and outer tubes is 0.2-8.0 m/s. [0053]
The hydrocarbon feedstocks fed to the inner tube and the annular space between the inner and outer tubes may be selected from a group consisting of: gaseous hydrocarbon, refinery gas, primary processing gasoline fraction, secondary processing gasoline fraction, primary processing diesel oil fraction, secondary processing diesel oil fraction, straight run gas oil, coker gas oil, deasphalted oil, hydrofined oil, hydrocracking tail oil, vacuum residuum and a mixture thereof The hydrocarbon feedstock fed to the inner tube is preferably selected from a group consisting of straight run gas oil, coker gas oil, deasphalted oil, hydrofined oil, hydrocracking tail oil, vacuum residuum, atmospheric residuum and a mixture thereof. The hydrocarbon feedstock fed to the annular space between the inner and outer tubes is preferably selected from a group consisting of gaseous hydrocarbon, refinery gas, primary processing gasoline fraction, secondary processing gasoline fraction, primary processing diesel oil fraction, secondary processing diesel oil fraction and a mixture thereof. The reaction conditions of the hydrocarbon feedstock in the inner tube are as follows a reaction temperature of 460-580° C.; preferably 480-550° C., a reaction pressure of 0.1-0 6 MPa, preferably 0.2-0.4 MPa; a catalyst-oil ratio (weight ratio of catalyst to feedstock) of 3-15, preferably 4-10; an oil-gas residence time of 1.0-10 seconds in the inner tube, preferably 1.5-5.0 seconds, a catalyst temperature of 620-720° C. before contacting the hydrocarbon feedstock, preferably 650-700° C.; and an atomization steam amount of 1-20% by weight (based on the hydrocarbon feedstock), preferably 2-15% by weight. [0054]
Reaction conditions of the hydrocarbon feedstock in the annular space between the inner and outer tubes are as follows: a reaction temperature of 300-680° C., preferably 400-600° C., a reaction pressure of 0.1-0.6 MPa, preferably 0.2-0.4 MPa; a catalyst-oil ratio of 2-30, preferably 4-20, an oil-as residence time of 0.5-20 seconds, preferably 1-15 seconds; and an atomization steam amount of 1-20% by weight (based on the hydrocarbon feedstock), preferably 1-15% by weight. The reaction conditions of the hydrocarbon feedstock in the annular space between the inner and outer tubes can be further optimized in the light of the properties of the hydrocarbon feedstock and the requirement of the desired product. When liquefied petroleum gas or light olefins are taken as major object products, relatively severe reaction conditions can be adopted; for example, a reaction temperature of 530-680° C., a catalyst-oil ratio of 8-30, an oil-gas residence time of 5-20 seconds and the like. When gasoline and/or diesel oil are taken as major object products, relatively moderate operation conditions should be adopted, for example, a reaction temperature 300-540° C., a catalyst-oil ratio of 2-9, an oil-gas residence time of 1-5 seconds, and the like. [0055]
The reaction stream both in the inner tube and in the annular space between the inner and outer tubes is mixed at the inlet of the confluence, and then continues to react in the confluence tube, wherein the reaction time is 0.1-3.0 seconds, and the reaction pressure, temperature, catalyst-oil ratio and the like depend on the reaction conditions in the inner tube and the annular space. Generally, the reaction temperature is 450-600° C., the catalyst-oil ratio is 4-12, the reaction pressure is 0.1-0.6 MPa, the weight ratio of steam to the hydrocarbon feedstock is 0.01-0.15:1. [0056]
The processes provided in the present invention are farther illustrated as follows in combination with the Figures, but not limited. [0057]
As shown in FIG. 6, the catalyst is fed to the bottom of the tube-in-tube riser reactor having single conduit for feeding catalyst via a [0058] catalyst inlet conduit 1, and flows upwards under an action of a pre-lifting media. A part of the catalyst (for example, 20-80 wt. % of the catalyst) flows into the inner tube 2 and the remaining part of the catalyst enters the annular space between the inner tube 2 and outer tube 3, both of them continue to flow upwards under the action of the pre-lifting media. The hydrocarbon feedstocks are fed via nozzles 8 and 9 to the inner tube and the annular space between the inner and outer tubes respectively to contact the catalyst, and both the reaction streams continue to flow upward along vessel wall. The reaction streams both from the inner tube and from the annular space between the inner and outer tubes flow together at the inlet of the confluence tube 4, then the mixed stream enters to the disengager 12 via the confluence tube and a rapid gas/solid separation apparatus, where a hydrocarbon product stream is separated from a spent catalyst; then the separated hydrocarbon product stream enters the subsequent separation system 14 to be separated further into different products. The spent catalyst falls down into the stripper 13, where the reaction oil-gas entrained by the catalyst is stripped off by the action of steam; the stripped catalyst is introduced to the regenerator 15 to burn off coke with air. The regenerated catalyst is recycled to the reactor for reuse.
As shown in FIG. 7, the catalyst is fed to the bottom of the tube-in-tube riser reactor having single conduit for feeding catalyst via the [0059] catalyst inlet conduit 1, and flows upward under the action of a pre-lifting media. A part of the catalyst (for example, 20-80 wt. % of the catalyst) flows into the inner tube 2, and the remaining part of the catalyst enters the annular space between the inner tube 2 and outer tube 3, both of them continue to flow upwards under the action of the pre-lifting media. The hydrocarbon feedstocks are fed via nozzles 8 and 9 into the inner tube and the annular space between the inner and outer tubes of the reactor respectively to contact the catalyst, and the reaction streams flow upwards continuously along vessel wall. The reaction streams both from the inner tube and from the annular space between the inner and outer tubes flow together at the inlet of the confluence tube 4, and the mixed stream enters the disengager 12 via the confluence tube and a gas/solid rapid separation apparatus, where a hydrocarbon product stream is separated from a spent catalyst. The separated hydrocarbon product stream enters the subsequent separation system 14 to be further separated into various products, and the spent catalyst falls down into the stripper 13 where the reaction oil-gas entrained by the catalyst is stripped off under the action of steam. The stripped catalyst is fed to the regenerator 15 to burn off coke with air, and then the regenerated catalyst is introduced to the blending vessel 16 to mix with the spent catalyst and/or semi-regenerated catalyst. The mixed catalyst is recycled to the reactor for reuse.
As show in FIG. 8, the regenerated catalyst is fed via [0060] catalyst inlet conduits 21 and 22 to the inner tube 35 and the annular space between the inner tube 35 and outer tube 36 in the tube-in-tube riser reactor having two conduits for feeding catalyst respectively, and flows upward under the action of pre-lifting media. The hydrocarbon feedstock is fed via nozzles 32 and 34 to the inner tube and the annular space between the inner and outer tubes respectively to contact the catalyst, and the reaction stream flows upward along vessel wall. The reaction streams both from the inner tube and from the annular space between the inner and outer tubes flow together at the inlet of the confluence tube 38, and the mixed stream enters the disengager 12 via the confluence tube and a rapid gas/solid separation apparatus. In the disengager, a hydrocarbon product stream is separated from a spent catalyst. The separated hydrocarbon product stream enters the subsequent separation system 14 to be further separated into various products The spent catalyst falls down into the stripper 13 where the reaction oil-gas entrained by the catalyst is stripped off under the action of steam, The stripped catalyst is fed to the regenerator 15 to burn coke off with air, and then the regenerated catalyst is recycled into the reactor for reuse.
As shown in FIG. 9, the regenerated catalyst is fed to the bottom of the [0061] inner tube 35 via catalyst inlet conduit 21 in the tube-in-tube riser reactor and flows upward under the action of a pre-lifting media A part of spent catalyst flows via the catalyst inlet conduit 22 into the annular space between the inner tube 35 and outer tube 36 and flows upward under the action of the pre-lifting media. The hydrocarbon feedstock is fed via nozzles 32 and 34 to the inner tube and the annular space between the inner and outer tubes respectively to contact the catalyst, and the reaction streams flow upward along vessel wall. Reaction streams both from the inner tube and from the annular space between the inner and outer tubes flow together at the inlet of the confluence tube 38, and the mixed stream enters the disengager 12 via the confluence tube and a rapid gas/solid separation apparatus. In the disengager, a hydrocarbon product stream is separated from a spent catalyst. Then the separated hydrocarbon product stream enters the subsequent separation system 14 to be further separated into various products, and the spent catalyst falls down into the stripper 13 where the reaction oil-gas entrained by catalyst is stripped off under the action of steam. The stripped spent catalyst is fed partially to the regenerator 15 to burn coke off with air, and the regenerated catalyst is recycled to the bottom of the inner tube for reuse The remaining part of the spent catalyst is recycled directly to the bottom of the annular space between the inner and outer tubes for reuse without regeneration.
According to the processes provided in the present invention, in addition to the flowcharts as shown in FIGS. [0062] 6-9 mentioned above, the regenerated catalyst, spent catalyst and semi-regenerated catalyst, or a mixture of any two thereof to be fed to the reactor can be fed via the catalyst inlet conduit to the bottom of the tube-in-tube riser reactor after being cooled via a catalyst cooler.
The reactor, catalyst, feedstock and operation conditions used in the relay cracking process of petroleum hydrocarbons provided in the present invention are substantially similar to those in the catalytic cracking process of petroleum hydrocarbons using tube-in-tube riser reactor aforementioned. The difference between them lies in that, the hydrocarbon feedstock is fed only to the inner tube of the tube-in-tube riser reactor but not to the annular space between the inner tube and outer tube. Thus, the relay cracking process of petroleum hydrocarbons is different from the catalytic cracking process of petroleum hydrocarbon using tube-in-tube riser reactor aforementioned in the aspects of feedstock and operation conditions. [0063]
The hydrocarbon feedstock fed into the inner tube is selected from a group consisting of gaseous hydrocarbon, refinery gas, primary processing gasoline fraction, secondary processing gasoline fraction, primary processing diesel oil fraction, secondary processing diesel oil fraction, straight run gas oil, coker gas oil, deasphalted oil, hydrofined oil, hydrocracking tail oil, vacuum gas oil, vacuum residuum, atmospheric residuum and a mixture thereof. Therein, it is preferred to select from the group consisting of straight run gas oil, coker gas oil, deasphalted oil, hydrofined oil, hydrocracking tail oil, vacuum gas oil, vacuum residuum, atmospheric residuum and a mixture thereof. [0064]
The reaction of the hydrocarbon feedstock in the inner tube is carried out under conditions as follows a reaction temperature of 480-700° C., preferably 500-680° C.; a reaction pressure of 0.1-0.6 MPa, preferably 0.2-0.4 MPa; a catalyst-oil ratio of 3-30, preferably 4-25; an oil-gas residence time of 1.0-10 seconds in the inner tube, preferably 1.5-5.0 seconds; a catalyst temperature of 620-800° C. before contacting hydrocarbon feedstock, preferably 640-750° C. and an atomization steam amount of 1-45% by weight (based on feedstock), preferably 2-35% by weight. [0065]
Reaction conditions of the hydrocarbons in the confluence tube are as follows: a reaction temperature of 490-720° C., preferably 500-700° C.; a reaction pressure of 0.1-0.6 MPa, preferably 0.2-0.4 MPa; a catalyst-oil ratio of 4-40, preferably 5-30; an oil-gas residence time of 0.5-10 seconds in the confluence tube, preferably 1.0-5.0 seconds; and a steam-oil ratio of 3-45% by weight, preferably 5-35% by weight. The steam-oil ratio means the weight ratio of steam to hydrocarbon feedstock. [0066]
The relay cracking process of petroleum hydrocarbons is further illustrated as follows in combination with the Figures, but is not limited. [0067]
As shown in FIG. 10, the regenerated catalyst is fed via the [0068] catalyst inlet conduit 1 to the bottom of the tube-in-tube riser reactor having single conduit for feeding catalyst, and flows upward under the action of a pre-lifting media; 20-80% by weight of catalyst flows into the inner tube 2, the remaining part of catalyst enters the annular space between the inner tube 2 and outer tube 3, and both the parts of the catalyst continue to flow upward under the action of the pre-lifting media. The hydrocarbon feedstock is fed to the inner tube of the reactor via nozzle 8 to contact the catalyst. And the reaction stream flows upward continuously along vessel wall. The reaction stream from the inner tube flows together at the inlet of the confluence tube 4 with the regenerated catalyst (stream) from the annular space between the inner and outer tubes, and then the mixed reaction stream enters the disengager 12 via the confluence tube and a rapid gas/solid separation apparatus, where a hydrocarbon product stream is separated from a spent catalyst. The separated hydrocarbon product stream enters the subsequent separation system 14 to be further separated into various products, and the spent catalyst falls down into the stripper 13 where the reaction oil-gas entrained by catalyst is striped off under the action of steam. The stripped catalyst is fed to the regenerator 15 to burn off coke with air, and the regenerated catalyst is recycled to the reactor for reuse.
As shown in FIG. 11, the regenerated catalyst is fed via the [0069] catalyst inlet conduits 21 and 22 to the inner tube 35 and the annular space between the inner tube 35 arid outer tube 36 of the tube-in-tube riser reactor having two conduits for feeding catalyst respectively, and flows upward under the action of a pre-lifting media The hydrocarbon feedstock is fed to the inner tube of the reactor via the nozzle 32 to contact the catalyst, and the reaction stream flows upward along vessel wall The reaction stream from the inner tube flows together at the inlet of the confluence tube 38 with the regenerated catalyst (stream) from the annular space between the inner tube 35 and outer tube 36 Then the mixed stream enters the disengager 12 via the confluence tube and a rapid gas/solid separation apparatus In the disengager, the hydrocarbon product stream and the spent catalyst are separated. The separated hydrocarbon product stream enters the subsequent separation system 14 to be further separated into various products. The spent catalyst falls down into the stripper 13 where the reaction oil-gas entrained by catalyst is stripped off under the action of steam. The stripped catalyst is introduced to the regenerator 15 to burn off coke with air, and the regenerated catalyst is recycled to the reactor for reuse.
In addition to those shown in FIGS. 10 and 11 aforementioned, in the relay cracking process of petroleum hydrocarbon, the regenerated catalyst to be fed to the reactor can be fed to the bottom of the tube-in-tube riser reactor after being cooled via a catalyst cooler. [0070]
The following Examples farther illustrate but do not intend to limit the processes provided in the present invention. [0071]
The catalyst used in the Examples is a commercial product from Catalyst Factory, Lan Zhou Petrochemical Corp., Trademark LV-23, of which major properties are shown in Table 1. The hydrocarbon feedstock used in the Examples is DaQing VGO blended with 30% by weight of VR, of which properties are shown in Table 2. The testing apparatus used in the Examples is a pilot FCC unit adopted tube-in-tube riser reactor. [0072]

EXAMPLE 1

The present example shows test results obtained by using the process provided in the present invention when light oil is produced as main object product. [0073]
Main steps of the test are as follows The feedstock shown in Table 1 and the recycled oil of the present unit were mixed, the mixed oil was heated via a preheating furnace and fed to the inner tube of the tube-in-tube riser reactor to contact the regenerated catalyst flowing from the regenerator and lifted with a pre-lifting media. The cracking gas produced in the present unit was fed to the annular space between the inner and outer tubes to contact the regenerated catalyst therein. The reaction stream in the inner tube and the reaction stream in the annular space moved up along vessel wall respectively, then mixed with each other in the confluence tube to continue the reaction, and then entered a disengager. In the disengager, the hydrocarbon product stream was separated from the spent catalyst, and then entered a subsequent fractionation system via an oil-gas pipeline to be further fractionated into various products. The products were metered and analyzed respectively. Spent catalyst was stripped with steam, then introduced into a regenerator to burn off coke with air. The regenerated catalyst was recycled to the reactor for reuse. [0074]
Main operation conditions, product distribution and main product properties are shown in Tables 3, 4 and 5 respectively. It can be seen from Tables 4 and 5 that when light oil is produced as main object product, the yield of gasoline+diesel oil up to 77.80% by weight and a total yield of liquid products can amount to 89.96% by weight with a lower yield of dry as and coke. [0075]

EXAMPLE 2

The present example shows test results obtained by using the process provided in the present invention when a liquefied petroleum gas is produced as main object product. [0076]
Main steps of the test are as follows The feedstock shown in Table 1 was heated via a preheating furnace, then fed to the inner tube of the tube-in-tube riser reactor to contact the regenerated catalyst flowing from the regenerator and lifted with a pre-lifting media. The light oil (gasoline diesel oil) with a distillation range less than 350° C. produced in the present unit was fed into the annular space between the inner and outer tubes to contact the regenerated catalyst therein. Both the reaction stream from the inner tube and the reaction stream from the annular space moved up along vessel wall respectively, mixed with each other in the confluence tube to continue the reaction, and then entered a disengager. In the disengager, the hydrocarbon product stream was separated from the spent catalyst, then entered a subsequent fractionation system via an oil-gas pipeline to be further fractionated into various products The products were metered and analyzed respectively. The spent Do catalyst was stripped with steam, and then introduced to a regenerator to burn off coke with air. The regenerated catalyst was recycled to the reactor for reuse. [0077]
The main operation conditions are shown in Table 3, the product distribution is shown in Table 4, and the main product properties are shown in Table 5 It can be seen from Tables 4 and 5 that when liquefied petroleum gas is produced as main object product, the yield of liquefied petroleum gas can amount to 29.46% by weight and the total yield of liquid products can amount to 86.65% by weight with a lower yield of dry gas and coke. [0078]

EXAMPLE 3

The present example shows test results obtained by using the process provided in the present invention when gasoline is produced as main object product. [0079]
The main steps of the test are as follows. The feedstock shown in Table 1 and the recycled oil produced in the present unit were mixed, then the mixed oil was heated via a preheating furnace and fed to the inner tube of the tube-in-tube riser reactor to contact the regenerated catalyst flowing from the regenerator and lifted with a pre-lifting media. Coker gasoline (with a density of 0.7316 g/cm[0080] ³, RON=58.8, MON=65.4, an olefin content of 37.49% by weight) was fed to the annular space between the inner and outer tubes to contact the regenerated catalyst therein. Both the reaction stream in the inner tube and the reaction stream in the annular space moved up along vessel wall respectively, mixed with each other in the confluence tube to continue the reaction, and then entered a disengager. In the disengager, the hydrocarbon product stream was separated from the spent catalyst, and then entered a subsequent fractionation system via an oil-gas pipeline to be further fractionated into various products The products were metered and analyzed respectively. The spent catalyst was stripped with steam, then introduced to a regenerator to burn off coke with air. The regenerated catalyst was recycled to the reactor for reuse.
Main operation conditions, product distribution and main product properties are shown in Tables 3, 4 and 5 respectively. It can be seen from Tables 4 and 5 that when gasoline is produced as main object product, the gasoline yield can amount to 59.49% by weight, the yield of light oil is 78.14% by weight, and total yield of liquid products amount to 90.00% by weight with a lower yield of dry as and coke meanwhile, a good quality of gasoline product is obtained. [0081]

EXAMPLE 4

The present example shows test results obtained by using the process provided in the present invention when diesel oil is produced as main object product. [0082]
Main steps of test are as follows. As shown in FIG. 7, the feedstock shown in Table I was mixed with the recycled oil of the present unit, the mixed oil was heated via a preheating furnace and fed to the inner tube of the tube-in-tube riser reactor to contact the mixed catalyst from [0083] catalyst blending vessel 16 in which the regenerated catalyst and the spent catalyst were mixed. The coker diesel oil (with a density of 0.8520 g/cm⁵, a sulfur content of 8225 ppm, and a nitrogen content of 5018 ppm, and a cetane number of 47) was fed to the annular space between the inner and outer tubes to contact the mixed catalyst therein. Both the reaction stream in the inner tube and the reaction stream in the annular space moved up along vessel wall respectively, then mixed with each other in the confluence tube to continue the reaction, and then entered a disengager In the disengager, the hydrocarbon product stream was separated from the spent catalyst, and then entered a subsequent fractionation system via an oil-gas pipeline to be further fractionated into various products. The products were metered and analyzed respectively. The spent catalyst was stripped with steam, then part of the stripped catalyst was introduced into a regenerator to burn off coke with air, and the remaining part of the spent catalyst was fed directly into the catalyst blending vessel 16 to mix with the regenerated catalyst in a blending ratio of the regenerated catalyst to the spent catalyst of 2:1. Then the mixed catalyst was recycled to the reactor for reuse.
Main operation conditions, product distribution and main product properties are shown in Tables 3, 4 and 5 respectively. It can be seen from Tables 4 and 5 that when diesel oil is produced as main object product, the yield of diesel oil can amount to 35.13% by weight and the total yield of liquid products can amount to 89.93% by weight with a lower yield of dry gas and coke. [0084]

EXAMPLE 5

The present example shows test results obtained by using the process provided in the present invention when liquefied petroleum gas and diesel oil are produced as main object products. [0085]
The main steps of test are as follows. The feedstock shown in Table 1 was mixed with the recycled oil from the present unit, and the mixed oil was heated via a preheating furnace and then fed to the inner tube of the tube-in-tube riser reactor to contact the regenerated catalyst flowing from regenerator and lifted with a pre-lifting media. The gasoline fraction produced in the present unit was fed to the annular space between the inner and outer tubes to contact the regenerated catalyst therein. Both the reaction stream in the inner tube and the reaction stream in the annular space moved up along vessel wall respectively, mixed each other in the confluence tube to continue the reaction, and then entered a disengager. In the disengager, the hydrocarbon product stream was separated from the spent catalyst, and entered a subsequent fractionation system via an oil-gas pipeline to be further fractionated into various products. The products were metered and analyzed respectively. The spent catalyst was stripped with steam, and then introduced to a regenerator to burn off coke with air. The regenerated catalyst was recycled to the reactor for reuse. [0086]
Main operation conditions, product distribution and main product properties are shown in Tables 3, 4 and 5 respectively. It can be seen from Tables 4 and 5 that when liquefied petroleum gas and diesel oil are produced as main object products, the yield of liquefied petroleum gas can amount to 19.08% by weight, the yield of diesel oil can amount to 31.46% by weight, and the total yield of liquid products can amount to 88.80% by weight with a lower yield of dry gas and coke. [0087]
In the following examples, the relay cracking process will be further illustrated but not limited. The catalysts used in the examples below were commercial products with trademarks respectively of RMG, CIP-1 and CFP manufactured by China's QiLu Petrochemical Catalyst Factory, and the three kinds of catalysts were hydrothermally aged, of which the main properties are shown in Table 6. In the examples, the hydrocarbon feedstock used is DaQing VGO blended with 30% by weight of VR, of which properties are shown in Table 1, and the testing apparatus used in the Examples below was a pilot FCC unit adopted tube-in-tube riser reactor. [0088]

EXAMPLE 6

The present example illustrates that higher yields of liquefied petroleum gas, gasoline and diesel oil can be obtained by using the relay cracking process provided in the present invention with heavy petroleum hydrocarbon as feedstock. [0089]
The steps of test are mainly as follows. As shown in FIG. 10, the regenerated catalyst was introduced into the bottom of the tube-in-tube riser reactor via the [0090] catalyst inlet conduit 1, and flew upward under the action of a pre-lifting vapor with 70% by weight of the catalyst flew into the inner tube 2 and the remaining 30% by weight of catalyst entered into the annular space between the inner tube 2 and outer tube 3 to continue flowing upward under the action of pre-lifting vapor. The hydrocarbon feedstock as shown in Table 1 was preheated, and fed to the inner tube via the nozzles 8 to contact the catalyst; the reaction stream flew upward continuously along vessel wall. The mixture of the reaction oil-gas and catalyst from the inner tube flew together in the confluence tube 4 with the regenerated catalyst stream from the annular space between the inner tube 2 and outer tube 3, and the mixed stream flew through the confluence tube and a rapid gas/solid separation apparatus into the disengager 12, where the hydrocarbon product stream was separated from the spent catalyst. The separated hydrocarbon product stream entered the subsequent separation system 14 to be further separated into various products. The spent catalyst fell down into the stripper 13 where the reaction oil-gas entrained by catalyst was stripped off under the action of steam and the stripped catalyst was introduced to the regenerator 15 to burn coke off with air. The regenerated catalyst was recycled to the reactor for reuse.
Both main operation conditions and product distribution are shown in Table 7. From Table 7, it can be seen that the present invention has a total liquid yield of light hydrocarbons (liquefied petroleum gas+gasoline+diesel oil) of 86.62% by weight with a lower yield of dry gas and coke. [0091]

EXAMPLE 7

The present example illustrates that a higher yield of light olefins such as propylene and the like can be obtained by using the relay cracking process of the present invention when a heavy petroleum hydrocarbon is used as feedstock. [0092]
Main steps of test are as follows. As shown in FIG. 11, the regenerated catalyst was fed via the [0093] catalyst inlet conduits 21 and 22 to the inner tube 35 and the annular space between the inner tube 35 and the outer tube 36 of the tube-in-tube riser reactor having two conduits for feeding catalyst respectively, and both parts of the regenerated catalyst flew upward under the action of pre-lifting media. The catalyst fed to the inner tube was 40% by weight of the total catalyst weight. The catalyst fed to the annular space was 60% by weight of the total catalyst weight. The hydrocarbon feedstock was fed to the inner tube of the reactor via the nozzle 32 to contact the catalyst, and the reaction stream flew upward along vessel wall. The reaction stream from the inner tube flew together in the confluence tube 38 with the regenerated catalyst stream from the annular space between the inner tube 35 and the outer tube 36, and then the mixed stream entered the disengager 12 via a confluence tube and a rapid gas/solid separation apparatus. In the disengager, a hydrocarbon product stream is separated from a spent catalyst. The separated hydrocarbon product stream entered the subsequent separation system 14 to be further separated into various products. The spent catalyst fell down into the stripper 13 where the reaction oil-gas entrained by catalyst was stripped off under the action of steam. The stripped catalyst was introduced to the regenerator 15 to burn coke off with air. The regenerated catalyst was recycled to the reactor for reuse.
Both main operation conditions and product distribution are shown in Table 7. From Table 7, it can be seen that when light olefins consisting mainly of propylene is produced as main object product, the yields of ethylene, propylene and butylene can amount to 4.79%, 24,01% and 15.28% by weight respectively with a lower yield of dry gas and coke. [0094]

EXAMPLE 8

This example illustrates that higher yield of light olefins such as ethylene and the like can be obtained by using the relay cracking process of the present invention with a heavy petroleum hydrocarbon as feedstock. [0095]
Main steps of test are as follows. As shown in FIG. 11, the regenerated catalyst was fed via the [0096] catalyst inlet conduits 21 and 22 to the inner tube 35 and the annular space between the inner tube 35 and the outer tube 36 of the tube-in-tube riser reactor having two conduits for feeding catalyst respectively, and both parts of the regenerated catalyst flew upward under the action of a pre-lifting media. The catalyst fed to the inner tube was in an amount of 50% of the total catalyst weight, and the catalyst fed to the annular space was also in an amount of 50% of the total catalyst weight. The hydrocarbon feedstock was fed to the inner tube of the reactor via the nozzle 32 to contact the catalyst, and the reaction stream flew upward along vessel wall The reaction stream from the inner tube flew together at the confluence tube inlet 38 with the regenerated catalyst stream from the annular space between the inner tube 35 and the outer tube 36, and then entered the disengager 12 via the confluence tube and a rapid gas/solid separation apparatus. In the disengager, the hydrocarbon product stream was separated from the spent catalyst. The separated hydrocarbon product stream flew into the subsequent separation system 14 to be further separated into various products. The spent catalyst fell down into the stripper 13 where the reaction oil-gas entrained by catalyst was stripped off under the action of steam. The stripped catalyst was introduced to the regenerator 15 to burn coke off with air. The regenerated catalyst was recycled to the reactor for reuse.
Both main operation conditions and product distribution are shown in Table 7. From Table 7, it can be seen that, the yields of ethylene, propylene and butylene can amount to 27.63%, 17.62% and 5.06% by weight, respectively, when light olefins consisting mainly of ethylene is produced as main object product by using the process of the present invention. [0097]

COMPARATIVE EXAMPLE

In the comparative example, a conventional riser reactor was used, and the same feedstock, catalyst and operation conditions were used as those used in Example 6. [0098]
Main steps of test are as follows. The preheated feedstock was fed to the riser reactor to contact the regenerated catalyst flowing from the regenerator and lifted with pre-lifting media. The formed reaction stream flew into a disengager via a riser; in the disengager, the hydrocarbon product stream was separated from the spent catalyst. Then, the separated hydrocarbon product stream entered into a subsequent fractionation system via oil-gas pipeline to be further fractionated into various products. The products were metered and analyzed respectively. The spent catalyst was stripped with steam, then introduced to a regenerator to burn off coke with air The regenerated catalyst was recycled into to the reactor for reuse. [0099]

Both main operation conditions and product distribution are shown in Table 7. Comparing Example 6 with the comparative example, it can be seen that the process of the present invention has a relatively strong ability to convert heavy oils and desired product selectivity.

	TABLE 1


		DaQing VGO blended
	Feedstock	with 30% of VR

	Density (20° C.), g/cm³	0.8881
	Refractive index (70° C.)	1.4784
	Kinematic viscosity, mm²/s, 80° C.	31.88
	100° C.	18.09
	Freezing point, ° C.	>50
	Aniline point, ° C.	112.9
	Conradson carbon value wt. %	2.7
	Four components, wt. %
	Saturated hydrocarbon	62.1
	Aromatics	25.2
	Resin	12.6
	Asphaltene	0.1
	Element composition, wt. %
	C	85.74
	H	13.01
	S	0.13
	N	0.20
	Ni, ppm	3.0
	Distillation range, ° C.
	Initial	339
	5%	388
	10%	421
	30%	473
	50%	526
	Uop K	12.7

	TABLE 2


	Catalyst	LV-23

	Chemical composition, wt. %
	Al₂O₃	51.2
	Na₂O	0.32
	RE₂O₃	2.0
	Physical Properties
	Specific surface area, m²/g	228
	Pore volume, ml/g	0.39
	sulk density, g/cm³	0.70
	Attrition index % h⁻¹	1.7
	Screen composition, v %
	0-20 μm	3.2
	0-40 μm	19.2
	0-80 μm	68.5
	0-110 μm	81.8
	0-149 μm	96.3
	Mean particle diameter, μm	66.8
	Aging condition in pilot unit	800° C./15 h/100% steam
	MA	61

TABLE 3


Example No.	1	2	3	4	5

object product	Gasoline + diesel	Liquefied petroleum	Gasoline	Diesel	Liquefied petroleum
		gas			gas + diesel
Operation mode	Combination operation	Combination	Combination	Combination	Combination
		operation	operation	operation	operation
Inner tube:
Feedstock	Fresh oil + heavy	Fresh oil	Fresh oil + heavy	Fresh oil + heavy	Fresh oil + heavy
	cycled oil		cycled oil	cycled oil	cycled oil
Fresh heavy oil feed, g/h	936	1190	1047	873	944
Reaction temperature, ° C.	510	550	515	530	520
Catalyst-oil ratio, w/w	5.1	8.0	6.2	4.5	5.0
Reaction time, s	1.62	1.58	1.60	1.50	1.55
Atomization steam, wt. %	5.7	6.1	5.8	5.9	5.7
Cycled ratio of heavy oil	0.28	0	0.15	0.39	0.26
Annular space:
Feedstock	Cracking gas	<350° C.	Coker gasoline	Coker diesel oil	FCC gasoline
		FCC fraction oil
Light oil feed, g/h	56 l/h	143	157	137	236
Light oil ratio(to fresh heavy	7.8	12.0	15.0	15.7	25.0
oil), wt. %
Reaction temperature, ° C.	570	620	610	540	580
Catalyst-oil ratio, w/w	6.1	12.2	8.0	4.9	7.9
Reaction time, s	2.46	1.67	1.32	1.47	0.91
Atomization steam, wt. %	3.5	8.8	8.6	8.6	8.7
Confluence tube.
Reaction temperature, ° C.	490	530	510	475	500
Catalyst-oil ratio, w/w	5.5	9.4	7.2	5.0	6.6
Reaction time, s	0.60	0.64	0.51	0.53	0.41
Atomization steam, wt. %	4.9	8.1	6.2	6.3	6.3

TABLE 4


Example No	1	2	3	4	5

Production Scheme	Gasoline +	LPG	Gasoline	LCO	LPG +
	LCO				LCO
Product distribution,
wt. %
Dry gas	2.89	3.26	2.79	2.49	3.23
LPG	12.16	29.46	11.86	10.72	19.08
Gasoline	49.28	40.70	59.49	44.08	38.26
LCO	28.52	16.49	18.65	35.13	31.46
Slurry	0.00	3.03	0.00	0.00	0.00
Coke	7.15	7.06	7.21	7.53	7.97
In total	100.00	100.00	100.00	100.00	100.00
Conversion, wt. %	73.48	80.48	81.35	64.87	68.54
Gasoline + LCO,	77.80	57.19	78.14	79.21	69.72
wt. %
LPG + Gasoline +	89.96	86.65	90.00	89.93	88.80
LCO, + wt. %

TABLE 5


Example No	1	2	3	4	5

Gasoline properties
Density(20° C.), g/cm³	0.7268	0.7312	0.7212	0.7255	0.7341
Actual gum, mg/100 ml	3	3	4	2	3
Mercaptan sulfur, ppm	34	38	40	53	35
Diene value, g I₂/100 g	1.2	1.8	1.6	0.8	1.1
Induction period. min	583	520	530	623	489
S, ppm	134	142	234	158	152
N, ppm	43	46	63	53	41
Octane number(measured)
RON	90.1	92.9	90.2	89.2	91.3
MON	78.4	82.5	78.6	78.1	79.5
Alkane content, v %	35.1	34.0	38.2	35.7	42.0
Olefin content, v %	52.3	42.6	46.5	50.7	37.8
Aromatics content, v %	12.6	23.4	15.3	13.6	20.2
Diesel Properties
Density, g/cm³	0.8718	0.8805	0.8732	0.8679	0.8709
Freezing point, ° C.	−3	−10	−4	−3	−10
Cetane number	42.4	38.4	41.0	40.2	42.6
S, wt. %	0.22	0.27	0.30	0.31	0.23
N, wt. %	0.01	0.03	0.06	0.02	0.02
Initial boiling point, ° C.	209	212	211	203	204
90% point distillated, ° C.	325	329	331	333	324

TABLE 6


Catalyst name	RMG	CIP-1	CEP

Chemical composition, m %
Al₂O	42.2	52.0	46.3
Na₂O	0.19	0.09	0.04
Physical Properties
Specific surface area, m²/g	206	210	152
Pore volume, cm³/g	0.22	0.3	0.24
Bulk density, g/cm³	0.81	0.8	0.86
Attrition index, m %/h	1.4	1.6	0.91
Screen composition, v %
0-20 μm	2.6	5.2	3.6
20-40 μm	10.3	20.4	13.7
40-80 μm	51.6	61.8	42.9
>80 μm	35.5	12.6	39.8
100% steam aging conditions:
Aging temperature, ° C.	800	800	820
Aging time, h	17	17	23
Activity of aged catalyst*, %	69	43*	53*

TABLE 7


Example No.	6	Comp. Ex.	7	8

Catalyst	RMG	RMG	CIP	1	CEP
Main operation conditions:		Riser
Inner tube:
Feeding amount,, g/h	1108	1100	490	363
Preheating temperature of	290	290	290	290
feedstock, ° C.
Reaction temperature, ° C.	495	515	530	620
Catalyst-oil ratio, w/w	5.5	7.8	5.8	15.7
Reaction time, s	1.62	2.93	1.58	1.47
Atomization water, wt. %	5.8	6.0	20.1	30.0
Regeneration temperature, ° C.	663	672	710	845
Confluence tube:
Reaction temperature, ° C.	515	/	580	650
Catalyst-oil ratio, w/w	7.9	/	14.5	31.4
Reaction time, s	1.86	/	1.56	1.73
Steam-oil Ratio, wt. %	5.8	/	22.8	36.7
Product distribution, wt. %
Dry gas	3.03	4.29	9.50	41.15
LPG	31.69	32.46	46.03	26.08
Gasoline	44.38	40.46	26.15	16.12
LCO	10.55	11.02	8.42	4.73
Slurry	4.96	6.04	3.22	1.54
Coke	5.39	5.73	6.88	10.38
In total	100.00	100.00	100.00	100.00
Conversion, wt. %	84.49	82.94	88.36	93.73
LPG + Gasoline + LCO, wt. %	86.62	83.94	80.60	46.93
Ethylene, wt. %	1.24	1.68	4.79	27.63
Propylene, wt. %	11.42	11.29	24.01	17.62
Butylene, wt. %	13.07	12.84	15.28	5.06

The present application claims priority under 35 U.S.C. §119 of Chinese Patent Applications No. 01258820.2 filed on Aug. 29, 2001, No. 01141477.4, filed on Sep. 27, 2001, No. 01264042.5, filed on Sep. 27, 2001, No. 01134268.4, filed on Oct. 30, 2001, and No. 01134269.2, filed on Oct. 30, 2001. The disclosures of each of the foregoing applications are expressly incorporated by reference herein in their entirety. [0107]

Claims

1. A catalytic cracking process of petroleum hydrocarbons, comprising the following steps:

(1) feeding a catalyst from an inlet conduit to an inner tube and an annular space between inner and outer tubes of a tube-in-tube riser reactor, which flows upward under an action of pre-lifting media;

(3) the reaction streams both from the inner tube and from the annular space between the inner and outer tubes flow together at the inlet of a confluence tube, and then enter a separation apparatus via the confluence tube, where a hydrocarbon product stream is separated from a spent catalyst;

2. The process according to claim 1, wherein said hydrocarbon feedstocks fed into the inner tube and the annular space between the inner and outer tubes are selected from a group consisting of gaseous hydrocarbon, refinery gas, primary processing gasoline fraction, secondary processing Gasoline fraction, primary processing diesel oil fraction, secondary processing diesel oil fraction, straight run gas oil, coker gas oil, deasphalted oil, hydrofined oil, hydrocracking tail oil, vacuum gas oil, vacuum residuum, atmospheric residuum and a mixture thereof.

3. The process according to claim 2, wherein said hydrocarbon feedstock fed into the inner tube is selected from a group consisting of straight run gas oil, coker gas oil, deasphalted oil, hydrofined oil, hydrocracking tail oil, vacuum gas oil, vacuum residuum, atmospheric residuum and a mixture thereof, and the hydrocarbon feedstock fed to the annular space between the inner and outer tubes is selected from a group consisting of gaseous hydrocarbon, refinery gas, primary processing gasoline fraction, secondary processing gasoline fraction, primary processing diesel oil fraction, secondary processing diesel oil fraction and a mixture thereof.

4. The process according to claim 1, wherein said hydrocarbon feedstock is reacted in the inner tube under conditions as follows: a reaction temperature of 460-580° C., a reaction pressure of 0 1-0.6 MPa, a catalyst-oil ratio of 3-15, an oil-gas residence time of 1.0-10 seconds in the inner tube, a catalyst temperature of 620-720° C. before contacting the feedstock, and an atomization steam amount of 3-20% by weight.

5. The process according to claim 4, wherein said hydrocarbon feedstock is reacted in the inner tube under conditions as follows: a reaction temperature of 480-550° C., a reaction pressure of 0.2-0.4 MPa, a catalyst-oil ratio of 4-10, an oil-gas residence time of 1.5-5.0 seconds in the inner tube, a catalyst temperature of 650-700° C. before contacting the feedstock, and an atomization steam amount of 2-15% by weight.

6. The process according to claim 1, wherein said hydrocarbon feedstock is reacted in the annular space between the inner and outer tubes under conditions as follows: a reaction temperature of 300-680° C., a reaction pressure of 0.1-0.6 MPa, a catalyst-oil ratio of 2-30, an oil-gas residence time of 0.5-20 seconds, and an atomization steam amount of 1-20% by weight.

7. The process according to claim 6, wherein said hydrocarbon feedstock is reacted in the annular space between the inner and outer tubes under conditions as follows: a reaction temperature of 400-600° C., a reaction pressure of 0.2-0.4 MPa, a catalyst-oil ratio of 4-20, an oil-gas residence time of l-15 seconds, and an atomization steam amount of 1-15% by weight.

8. The process according to claim 1, wherein said catalyst fed to the inner tube and the annular space between the inner and outer tubes of the tube-in-tube riser reactor is selected from a group consisting of: a regenerated catalyst, a semi-regenerated catalyst, a spent catalyst and a mixture thereof, and the carbon content of the catalyst fed to the inner tube may be different from that of the catalyst fed to the annular space between the inner tube and outer tubes.

9 The process according to claim 1, wherein said tube-in-tube riser reactor is one having single conduit for feeding catalyst, which comprises mainly the following members: catalyst inlet conduit (1), inner tube (2), outer tube (3), confluence tube (4), pre-lifting distribution rings (5), (6) and (7) and feed nozzles (8) and (9); wherein inner tube (2) and outer tube (3) are coaxial, and the ratio of the cross-section area of the inner tube to the cross-section area of the annular space between the inner and outer tubes is 1:0.1-10; the lower end of the inner tube (2) is located at a place above the catalyst inlet, the inner tube has a length amounting to 10-70% of the total length of the reactor, one end of the confluence tube (4) is connected with the upper end of the outer tube (3), and the other end is connected with a gas/solid separation apparatus, the cross-section area ratio of the confluence tube (4) to the inner tube (2) is 1:0.2-0.8; the pre-lifting distribution rings (5), (6) and (7) are respectively located at the bottoms of the reactor, the inner tube and the outer tube.

10. The process according to claim 9, wherein said ratio of the cross-section area of the inner tube to the cross-section area of the annular space between the inner and outer tubes in said tube-in-tube riser reactor is 1:0.2-2.

11. The process according to claim 9, wherein said inner tube of the tube-in-tube riser reactor has a length amounting to 20-60% of the total length of the reactor.

12. The process according to claim 9, wherein said tube-in-tube riser reactor has a distance of 1-30 meters from the outlet end of the inner tube to the outlet end of the confluence tube

13. The process according to claim 9, characterized in that said tube-in-tube riser reactor is installed with 2-12 lines of shrouding wires or draw-bars between the inner tube and the outer tube.

14. The process according to claim 1, wherein said tube-in-tube riser reactor is one having two conduits for feeding catalyst, and comprises mainly the following members: catalyst inlet conduits (21) and (22), inner tube (35), outer tube (36), confluence tube (38), pre-lifting distribution rings (31) and (33), and feed nozzles (32) and (34), wherein the inner tube (35) and outer tube (36) are coaxial; the ratio of the cross-section area of the inner tube to the cross-section area of the annular space between the inner and outer tubes is 1:0.1-10; the catalyst inlet conduit (21) is connected with the lower end of the inner tube (35), the length of the inner tube is 10-70% of the total length of the reactor; the distance from the lower end of the outer tube (36) to the lower end of the inner tube (35) is 2-20% of the total length of the reactor; the catalyst inlet conduit (22) is connected with the lower end of the outer tube (36); one end of the confluence tube (3) is connected with the upper end of the outer tube (36) and the other end is connected with a gas/solid separation apparatus, the cross-section area ratio of the confluence tube (38) to the inner tube (35) is 1:0.2-0.8, the pre-lifting distribution rings (31) and (33) are located at the bottom of the inner tube and the bottom of the annular space between the inner tube and outer tube respectively, and feed nozzles (32) and (34) are located at the lower parts of the inner tube and the outer tube respectively.

15. The process according to claim 14, wherein said ratio of the cross-section area of the inner tube to the cross-section area of the annular space between the inner and outer tubes is 1.0.1-2 in said tube-in-tube riser reactor.

16. The process according to claim 14, wherein said tube-in-tube riser reactor has an inner tube with a length amounting to 15-50% of the total length of the riser reactor.

17. The process according to claim 14, wherein said distance from the lower end of the outer tube to the lower end of the inner tube in said tube-in-tube riser reactor is 5-15% of the total length of the riser reactor.

18. The process according to claim 14, wherein said cross-section area ratio of the confluence tube to the inner tube is 1:0.3-0.7 in said tube-in-tube riser reactor.

19. The process according to claim 14, wherein said tube-in-tube riser reactor is installed with 2-12 lines of shrouding wires or draw-bars between the inner tube and the outer tube.

20. The process according to claim 14, wherein said tube-in-tube riser reactor has a distance of 0.5-20 meters from the outlet end of the inner tube to the outlet end of the confluence tube.

21 A catalytic cracking process of petroleum hydrocarbons, comprising mainly the following steps:

(4) separating the hydrocarbon product stream further into various products including gasoline, diesel oil and liquefied petroleum gas; stripping and regenerating the spent catalyst, and recycling the regenerated catalyst into the reactor for reuse.

22. The process according to claim 21, wherein said catalyst fed to the tube-in-tube riser reactor via the catalyst inlet conduit is a regenerated catalyst or a cooled regenerated catalyst.

23. The process according to claim 21 or 22, wherein the active components of said catalyst are at least one kind of zeolite selected from a group consisting of Y-type or HY-type zeolite containing or not containing rare earth and/or phosphorous, ultra-stable Y-type zeolite containing or not containing rare earth and/or phosphorous, ZSM-5 family zeolite or high-silica zeolite having a pentasil structure, β-zeolite, ferrierite and a mixture thereof.

24. The process according to claim 21, wherein said catalyst contains 0.5-60% by weight of ZSM-5 family zeolite or other high-silica zeolite having a pentasil structure.

25. The process according to claim 21, wherein said hydrocarbon feedstock fed to the inner tube is selected from a group consisting of gaseous hydrocarbon, refinery gas, primary processing gasoline fraction, secondary processing gasoline fraction, primary processing diesel oil fraction, secondary processing diesel oil fraction, straight run gas oil, coker gas oil, deasphalted oil, hydrofined oil, hydrocracking tail oil, vacuum gas oil, vacuum residuum, atmospheric residuum and a mixture thereof.

26. The process according to claim 25, wherein said hydrocarbon feedstock fed to the inner tube is selected from the group consisting of straight run gas oil, coker gas oil, deasphalted oil, hydrofined oil, hydrocracking tail oil, vacuum gas oil, vacuum residuum, atmospheric residuum and a mixture thereof.

27. The process according to claim 21, wherein said hydrocarbon feedstock is reacted in the inner lube under conditions as follows, a reaction temperature of 480-700° C. a reaction pressure of 0.1-0.6 MPa, a catalyst-oil ratio of 3-30, an oil-gas residence time of 1.0-10 seconds in the inner tube, a catalyst temperature of 620-800° C. before contacting the feedstock, and an atomization steam amount of 1-45% by weight.

28. The process according to claim 27, wherein said hydrocarbon feedstock is reacted in the inner tube under conditions as follows: a reaction temperature of 500-680° C., a reaction pressure of 0.2-0.4 MPa, a catalyst-oil ratio of 4-25, an oil-gas residence time of 1.5-5.0 seconds in the inner tube, a catalyst temperature of 640-750° C. before contacting the feedstock, and an atomization steam amount of 2-35% by weight.

29. The process according to claim 21, wherein said reaction stream is reacted in the confluence tube under conditions as follows: a reaction temperature of 490-720° C., a reaction pressure of 0.1-0.6 MPa, a catalyst-oil ratio of 4-40, an oil-gas residence time of 0.5-10 seconds in the confluence tube, and a steam-oil ratio of 3-45% by weight.

30. The process according to claim 29, wherein said reaction stream is reacted in the confluence tube under conditions as follows: a reaction temperature of 500-700° C., a reaction pressure of 0.2-0.4 MPa, a catalyst-oil ratio of 5-30, an oil-gas residence time of 1.0-5 seconds in the confluence tube, and a steam-oil ratio of 5-35% by weight.

31. The process according to claim 21, wherein said tube-in-tube riser reactor is one having single conduit for feeding catalyst and comprises mainly the following members: catalyst inlet conduit (1), inner tube (2), outer tube (3), confluence tube (4), pre-lifting distribution rings (5), (6) and (7), and feed nozzles (8) and (9), wherein, the inner tube (2) and the outer tube (3) are coaxial, and the ratio of the cross-section area of the inner tube to the cross-section area of the annular space between the inner and outer tubes is in the range 1:0.1-10; the lower end of the inner tube (2) is located at a place above the catalyst inlet; the inner tube has a length amounting to 10-70% of the total length of the reactor; one end of the confluence tube (4) is connected with the upper end of the outer tube (3), and the other end is connected to the gas/solid separation apparatus, the cross-section area ratio of the confluence tube (4) to the inner tube (2) is 1:0.2-0.8; the pre-lifting distribution rings (5), (6) and (7) are located at the bottoms of the reactor, the inner tube and the outer tube respectively.

32. The process according to claim 31, wherein said tube-in-tube riser reactor has a ratio of the cross-section area of the inner lube to the cross-section area of the annular space between the inner and outer tubes in a range of 1:0.2-2.

33. The process according to claim 31, wherein said inner tube of the tube-in-tube riser reactor has a length amounting to 20-60% of the total length of the reactor.

34 The process according to claim 31, wherein said tube-in-tube riser reactor has a distance of 1-30 meters from the outlet end of the inner tube to the outlet end of the confluence tube.

35. The process according to claim 31, characterized in that said tube-in-tube riser reactor is installed with 2-12 lines of shrouding wires or draw-bars between the inner tube and outer tube.

36. The process according to claim 21, wherein said tube-in-tube riser reactor is one having two conduits for feeding catalyst and comprises mainly the following members: catalyst inlet conduits (21) and (22), inner tube (35), outer tube (36), confluence tube (38), pre-lifting distribution rings (31) and (33), and feed nozzles (32) and (34), wherein, the inner tube (35) and outer tube (36) are coaxial, the ratio of the cross-section area of the inner tube to the cross-section area of the annular space between the inner and outer tubes is 1:0.1-10; the catalyst inlet conduit (21) is connected with the lower end of the inner tube (35), the length of the inner tube is 10-70% of the total length of the reactor; the distance from the lower end of the outer tube (36) to the lower end of the inner tube (35) is 2-20% of the total length of the reactor, the catalyst inlet conduit (22) is connected with the lower end of the outer tube (36); one end of the confluence tube (38) is connected with the upper end of the outer tube (36), and the other end is connected with a gas/solid separation apparatus, the cross-section area ratio of the confluence tube (38) to the inner tube (35) is 1:0.2-0.8; the pre-lifting distribution rings (31) and (33) are located at the bottom of the inner tube and the bottom of the annular space between the inner and outer tubes respectively; feed nozzles (32) and (34) are located at the lower part of the inner tube and the lower part of the outer tube respectively.

37. The process according to claim 36, wherein said tube-in-tube riser reactor has a ratio of the cross-section area of the inner tube to the cross-section area of the annular space between the inner and outer tubes in a range of 1.0.1-2

38. The process according to claim 36, wherein said inner tube of said tube-in-tube riser reactor has a length amounting to 15-50% of the total length of the riser reactor.

39. The process according to claim 36, wherein said tube-in-tube riser reactor has a distance of 5-15% of the total length of the riser reactor from a lower end of the outer tube to the lower end of the inner tube.

40. The process according to claim 36, wherein said tube-in-tube riser reactor has an cross-section area ratio of the confluence tube to the inner tube in a range of 1:0.3-0.7.

41. The process according to claim 36, wherein said tube-in-tube riser reactor is installed with 2-12 lines of shrouding wires or draw-bars between the inner tube and outer tube.

42. The process according to claim 36, wherein said tube-in-tube riser reactor has a distance of 0.5-10 meters from the outlet end of the inner tube to the outlet end of the confluence tube.