US20020195373A1

US20020195373A1 - Heavy oil fluid catalytic cracking process

Info

Publication number: US20020195373A1
Application number: US09/883,465
Authority: US
Inventors: Takashi Ino; Toshiaki Okuhara; Halim Redhwi; Mohammad Abul-Hamayel; Abdullah Aitani; Abdulgader Maghrabi
Original assignee: Individual
Current assignee: Individual
Priority date: 2001-06-07
Filing date: 2001-06-07
Publication date: 2002-12-26

Abstract

A vacuum gas oil treated with hydrogen, with the hydrogen partial pressure higher than 80 kg/cm²G, is catalytically cracked in a fluid catalytic cracking apparatus having a regeneration zone, reaction zone, separation zone, and stripping zone, under conditions that a reaction zone outlet temperature is in the range of 550 to 630° C. and a contact time of hydrocarbons in the reaction zone is in the range of 0.01 to 1.0 sec. According to the fluid catalytic cracking process, a yield of light fraction olefins can be enhanced while a yield of coke can be lessened.

Description

BACKGROUND OF THE INVENTION

1. Field of the Invention

This invention relates to a process for catalytic cracking of a heavy fraction oil. More particularly, it relates to a fluid catalytic cracking (FCC) process which comprises cracking a vacuum gas oil to obtain light olefins such as ethylene, propylene, and butenes.

2. Description of the Prior Art

In a usual catalytic cracking technique, petroleum-derived hydrocarbons are catalytically cracked with a catalyst thereby to obtain gasoline as the main product, a small amount of LPG, a cracked gas oil and the like, and coke deposited on the catalyst is then burnt away with air to recycle the regenerated catalyst for reuse.

In recent years, however, there has been a tendency that a fluid catalytic cracking apparatus is utilized not as an apparatus for producing gasoline but as an apparatus for producing light fraction olefins for use as petrochemical materials. Such utilization of an original fluid catalytic cracking apparatus as an olefin-producing apparatus is economically advantageous, particularly to an oil refinery which is highly integrated with a petrochemical industry.

On the other hand, much attention has been paid to environmental problems, and therefore, regulation of the content of olefins and aromatics in gasoline for automobiles. In consequence, it can be anticipated that alkylate will be increasingly in demand as base materials for high-octane gasoline, rather than FCC gasoline and catalytically-reformed gasoline. Therefore, it will be necessary to increase the production of propylene and butenes, which are raw materials for alkylate production.

Methods for producing the light fraction olefins by the fluid catalytic cracking of heavy fraction oils include methods which comprise contacting a feed oil with a catalyst for a short time (U.S. Pat. Nos. 4,419,221, 3,074,878, and 5,462,652, and European Patent No. 315,179A), a method which comprises carrying out a cracking reaction at a high temperature (U.S. Pat. No. 4,980,053), and methods which comprise using pentasil type zeolites (U.S. Pat. No. 5,326,465 and Japanese Patent National Publication (Kohyo) No. Hei 7-506389 (506389/95)).

Even these known methods still cannot sufficiently produce light-fraction olefins selectively. For example, the high-temperature cracking reaction will result in concurrent thermal cracking of heavy fraction oils, thereby increasing the yield of dry gases from said oils; the short contact time of a feed oil with a catalyst will decrease the conversion of light-fraction olefins to light-fraction paraffins due to its inhibition of a hydrogen transfer reaction, but it will be unable to increase conversion of heavy-fraction oils to light-fraction oils; and, moreover, the use of pentasil-type zeolites will only enhance the yield of light-fraction hydrocarbons by excessively cracking the gasoline once it is produced. Therefore, it is difficult to produce light-fraction olefins in a high yield from heavy-fraction oils by using each of these known techniques alone.

SUMMARY OF THE INVENTION

An objective of this invention is to provide an improved process for the fluid catalytic cracking of a heavy-fraction oil, which can produce light-fraction olefins in a high yield while producing a lessened amount of dry gases such as gaseous hydrogen, methane, and ethane generated by the thermal cracking (thermocracking) of the heavy-fraction oil.

The present inventors, in an attempt to mainly increase the yield of light-fraction olefins, while inhibiting thermal cracking which will produce a large amount of dry gases, have intensively studied a process for the fluid catalytic cracking of a heavy-fraction oil at a high temperature. As a result, they have found that the objective can be achieved by contacting a vacuum gas oil treated with hydrogen, under specific conditions to be described below, with a catalyst at a suitable temperature. This invention has been achieved on the basis of these findings.

More particularly, the process for the fluid catalytic cracking of a heavy-fraction oil, according to this invention, comprises the step of contacting a vacuum gas oil treated with hydrogen, under the hydrogen partial pressure higher than 80 kg/cm ²G, with a catalyst in a fluid catalytic cracking apparatus having a regeneration zone, reaction zone separation zone, and stripping zone, and under conditions that a reaction zone outlet temperature is in the range of 550 to 630° C. and a contact time of hydrocarbons in the reaction zone is in the range of 0.01 to 1.0 sec.

DETAILED DESCRIPTION OF THE INVENTION

Apparatus and Process

The fluid catalytic cracking apparatus which can be used in this invention has a regeneration zone (a regenerator), a reaction zone (a reactor), a separation zone (a separator), and a stripping zone (a stripper).

The term “fluid catalytic cracking” referred to herein indicates that the above-described heavy-fraction oil, as the feed oil, is continuously brought into contact with a catalyst kept in a fluidized state under specific operating conditions, to crack the heavy-fraction oil, thereby producing light-fraction hydrocarbons, mainly comprising gasoline and light-fraction olefins. One type of reaction zone used in this invention is a up-flow type (so-called riser) reaction zone, wherein both catalyst particles and feed oil ascend through a pipe. Another type of reaction zone used in this invention is a down flow-type (so-called downer) reaction zone, wherein both catalyst particles and feed oil descend through a pipe. In case of a riser reactor, however, back-mixing of hydrocarbons takes place in the reactor, which brings a localized long residence time of hydrocarbons in the reactor, resulting in enhancement of thermal cracking. Particularly, as is the case in this invention, fluid catalytic cracking processes operated with higher reaction temperatures than ordinal fluid catalytic cracking processes, thermal cracking caused by back-mixing is significant. With an increase of the contribution of thermal cracking, undesirable dry gases increase, and desirable gasoline and light fraction olefins decrease. In this invention, the downer reactor is preferably adopted, to minimize thermal cracking.

BRIEF DESCRIPTION OF THE DRAWING

FIG. 1 is a schematic illustration of an example of a fluid catalytic cracking (FCC) apparatus. In this example, the fluid catalytic cracking apparatus has a down-flow type reaction zone.[0015]
This invention is more fully explained in the context of an FCC process. The drawing of this invention shows a typical FCC process arrangement. The description of this invention in the context of this specific process arrangement shown is not meant to limit it to the details disclosed therein. The principal components of the FCC arrangement shown in FIG. 1 consist of a reaction zone [0016] 1, a gas-solid separation zone 2, a stripping zone 3, a regeneration zone 4, a riser type regenerator 5, a catalyst hopper 6, and a mixing zone 7. The arrangement circulates catalyst and contacts feed in the manner hereinafter described.
An FCC feedstock such as heavy oil is charged to the mixing zone [0017] 7 through line 10, and there is mixed with the regenerated catalyst from the catalyst hopper 6. The mixture falls downward through the reaction zone 1, where the cracking reaction of heavy oil takes place under high reaction temperatures and at short contact time. Then, the mixture of spent catalyst and products, from the reaction zone 1, enters the gas-solid separation zone 2 located under the reaction zone 1. The spent catalyst is separated, in separation zone 2, from the cracked products and un-reacted feed oil. The catalyst is then sent to the upper portion of the stripping zone 3 through dip leg 9.
Hydrocarbon gases separated from most of the spent catalyst are sent to a secondary separator [0018] 8, where the rest of the spent catalyst is separated from the product gas. Hydrocarbon gases are then sent to a product recovery section. A tangential-type cyclone is preferred for use as the secondary separator 8.
Catalyst separated by the secondary separator [0019] 8 is directed to the stripping zone 3 where heavy hydrocarbons adsorbed on the catalyst are removed by a stripping gas introduced through line 11. Steam produced in a boiler, or an inert gas such as nitrogen, pressurized in a compressor, is preferably used as the stripping gas. As stripping conditions, a stripping temperature of 500 to 630° C. and a catalyst residence time of 1 to 10 minutes are preferred. Vapors of cracked products and un-reacted feed oil, stripped from the spent catalyst in the stripping zone 3, are withdrawn through line 12 located at the top of the stripping zone, together with the stripping gas. These gases are then sent to a product recovery section. The spent stripped catalyst is transferred to the regeneration zone through the line that has the first flow controller 13.
The superficial gas velocity in the [0020] stripping zone 3 is preferably maintained within the range of 0.05 to 0.4 m/s, in order to keep the fluidized bed in the stripping zone in a bubble phase. Since the gas velocity is relatively low within the bubble phase zone, the consumption of stripping gas can be minimized. Moreover, the range of operational pressure of the first flow controller 13 may be broad, during the bubble phase condition, due to the high bed density, and therefore the transportation of catalyst particles from the stripping zone 3 to the regeneration zone 4 is facilitated. Perforated trays or other internal structures can be used in the stripping zone 3 to improve stripping efficiency between the stripping gas and the catalyst.
The regeneration zone [0021] 4 is composed of a cone-shaped column connected at the top to a vertical line 5 (a riser-type regenerator). The tip angle of the cone is preferably in the range of 30° to 90°. The ratio of the cone height to the column diameter is preferably in the range of ½ to 2.
The spent catalyst is regenerated with a combustion gas (typically an oxygen-containing gas such as air) which is fed to the regeneration zone [0022] 4 through line 14. Regeneration is by completely burning, under fluidized conditions, the carbonaceous materials and heavy hydrocarbons which have been adsorbed on the spent catalyst. Catalyst regeneration temperature is normally in the range of 600° to 1000° C., preferably in the range of 650°to 750° C. Catalyst residence time in the regeneration zone 4 is in the range of 1 to 5 minutes, and the superficial gas velocity is preferably in the range of 0.4 to 1.2 m/s.
After regeneration of the spent catalyst in the regeneration zone [0023] 4, the regenerated catalyst in the upper portion of the turbulent-phase fluidized bed is transferred to a riser-type regenerator 5.
The ratio of the riser-type regenerator [0024] 5 diameter, to the diameter of the column located at the bottom region of the regeneration zone, is preferably in the range of ⅙ to ⅓. By adopting this diameter ratio, the superficial gas velocity in the regeneration zone 4 can be kept within 0.4 to 1.2 m/s, a velocity value which is appropriate to cause a turbulent fluidized bed. Moreover, the superficial gas velocity in the riser-type regenerator 5 can be kept within 4 to 12 m/s, a value which is appropriate for the upward transportation of the regenerated catalyst.
The regenerated catalyst from the riser-type regenerator [0025] 5 is carried to the catalyst hopper 6 located at the top of the riser type regenerator. The catalyst hopper 6, which functions as a gas-solid separator, where the flue gases that contain the by-products of coke combustion are separated from the regenerated catalyst and removed through cyclone 15.
The regenerated catalyst in catalyst hopper [0026] 6 is routed to the mixing zone 7 through a downer line equipped with a second flow controller 17. If necessary, a portion of the regenerated catalyst in the catalyst hopper 6 can be returned back to the regeneration zone 4 through a bypass line equipped with a third flow controller 16.
As described above, FCC catalyst circulates in the HS-FCC apparatus through a reaction zone [0027] 1, a gas-solid separation zone 2, a stripping zone 3, a regeneration zone 4, a riser-type regenerator 5, a catalyst hopper 6, and a mixing zone 7.
A mixture of the catalyst and the products obtained by the catalytic cracking of the heavy oil in contact with the catalyst in a reaction zone, is then forwarded into the separation zone, in which most of the catalyst is separated from the hydrocarbon gas. Next, the separated catalyst is forwarded to the stripping zone, in which most of heavy hydrocarbons adsorbed on the catalyst are removed from the catalyst particles. The catalyst, on which carbonaceous materials and a portion of heavy hydrocarbons are deposited, then is forwarded from the stripping zone to the regenerating zone. In the regenerating zone, the catalyst on which the carbonaceous materials and the like are deposited, is subjected to oxidation treatment, to decrease the amount of the deposits, thereby obtaining a regenerated catalyst. This regenerated catalyst is continuously recycled to the reaction zone. In one particular case, the cracked products are quenched just upstream of, or just downstream of, the separator, in order to restrict unnecessary further cracking or excessive cracking. [0028]
The “reaction zone outlet temperature” referred to in this invention means an outlet temperature of the reaction zone, and it is a temperature before separation of the cracked products from the catalyst, or, a temperature before quenching the cracked products, in case that they are quenched just upstream of the separator. In this invention, the reaction zone outlet temperature is in a range of 550 to 630° C., preferably 580 to 620° C. If the reaction zone outlet temperature is lower than 550° C., then the light-fraction olefins will not be obtained in a high yield, while if the reaction zone outlet temperature is higher than 630° C., then the thermal cracking of heavy-fraction oil feedstock will be significant, thereby undesirably increasing the amount of dry gases generated. [0029]
The contact time referred to herein means either the time between the start of contact of the feed oil with the regenerated catalyst, and the separation of the produced cracked products from the catalyst, or the time between the start of contact of the feed oil with the regenerated catalyst, and the quenching, in the configuration in which the obtained cracked products are quenched just upstream of the separation zone. In this invention, the contact time is in a range of 0.01 to 1.0 sec, preferably 0.05 to 0.8 sec, more preferably 0.1 to 0.6 sec. If the contact time is less than 0.01 sec, then the light fraction olefins will not be obtained in a high yield, because of the low conversion of the heavy-fraction oil, while if the contact time is longer than 1.0 sec, then the thermal cracking of heavy-fraction oil feed will be significant, thereby undesirably increasing the amount of dry gases generated. [0030]
In this invention, although operating conditions of the fluid catalytic cracking apparatus, except those described above, are not particularly restricted, the apparatus can be operated preferably at a reaction pressure of 1 to 3 kg/cm[0031] ²G, at a regeneration zone temperature of 650 to 750° C., and at a catalyst/oil ratio of 8 to 40 wt/wt, preferably 12 to 30 wt/wt. Said catalyst/oil ratio means the ratio of the amount of the catalyst recycled (ton/hr) to the feed rate (ton/hr).
Feed Oil (Feedstock or Charge Stock) [0032]
In the fluid catalytic cracking process in this invention, a vacuum gas oil (effluent from a vacuum distillation unit), which has been treated in a hydrotreating unit, is used as a feed oil. A hydrogen partial pressure within said hydrotreating unit is higher than 80 kg/cm[0033] ²G, preferably 145 kg/cm²G. By hydrotreating a vacuum gas oil under such high hydrogen partial pressure, nitrogen compounds which poison the catalyst are almost completely removed from the vacuum gas oil, and furthermore, aromatic compounds included in the vacuum gas oil is are hydrogenated to paraffins. Therefore, the crackability of the vacuum gas oil is much improved by the prior hydrotreatment. If the hydrogen partial pressure of said hydrotreating unit is less than 80 kg/cm²G, the light-fraction olefins will not be obtained in high yield, even though the hydrotreated vacuum gas oil is supplied as a feed oil to the fluid catalytic cracking apparatus operated under the conditions specified in this invention, because the extent of hydrogenation of aromatic compounds in the vacuum gas oil was insufficient. In this invention, operating conditions of the hydrotreating unit for vacuum gas oils are not particularly restricted,except as described above.
The hydrotreating unit can be operated preferably at a reaction temperature of 350 to 430° C., at an LHSV of 0.5 to 4.0, and at a H[0034] ₂/Oil ratio of 1,000 to 4,000 scf/bbl. Preferable properties of the vacuum gas oil used as a feed oil for the fluid catalytic cracking apparatus in this invention are as follows: hydrogen content is more than 13 wt %, more preferably more than 13.5 wt %; nitrogen content is less than 0.01 wt %, more preferably less than 0.005 wt %; sulfur content is less than 0.1 wt %; specific gravity at 15° C. is in the range of 0.80 to 0.85 g/cm³; and aniline point is in the range of 100 to 130° C. Preferable distillation properties of said vacuum gas oil are as follows; 10 vol % point is in a range of 250 to 470° C., more preferably in a range of 340 to 420° C.; 50 vol % point is in a range of 300 to 520° C., more preferably in a range of 400 to 500° C.; 90 vol % point is in a range of 340 to 560° C., more preferably in a range of 470 to 550° C.
The catalyst which is used in this invention contains the ultra stable Y-type zeolite which is the active component, and a matrix which is a substrate material for the zeolite. The matrices include clays such as kaolin, montmorilonite, halloysite and bentonite, and inorganic porous oxides such as alumina, silica, boria, chromia, magnesia, zirconia, titania and silica-alumina. [0035]
The content of the ultra stable Y-type zeolite in the catalyst used in this invention is preferably in a range of 5 to 50 wt %, more preferably 15 to 40 wt %. [0036]
The catalyst used in this invention may contain, in addition to the ultra stable Y-type zeolite, a crystalline aluminosilicate zeolite or silicoaluminophosphate (SAPO), each having smaller pores than the ultrastable Y-type zeolite. The aluminosilicate zeolites and SAPOs include ZSM-5, beta, omega, SAPO-5, SAPO-11 and SAPO-34. The zeolite or the SAPO may be contained in the catalyst particles containing the ultrastable Y-type zeolite, or may be contained in other catalyst particles. [0037]
The catalyst used in this invention preferably has a bulk density of 0.5 to 1.0 g/ml, an average particle diameter of 50 to 90 microns, a surface area of 50 to 350 m[0038] ²/g and a pore volume of 0.05 to 0.5 ml/g.

DESCRIPTION OF THE PREFERRED EMBODIMENTS

Next, this invention will be described with reference to the following examples and the like, but this invention should not be limited to these examples. [0039]

Example 1

Catalytic cracking of a heavy oil was carried out by the use of an adiabatic down flow type FCC pilot device. With regard to the scale of the device, the inventory (the amount of catalyst) was 5 kg and the feed oil rate was 1 kg/hr, and with regard to operation conditions, the reaction zone outlet temperature was 600° C., the reaction pressure was 1.0 kg/cm[0040] ²G, the catalyst/oil ratio was 15.5 wt/wt, the contact time was 0.4 seconds, and the regenerating zone catalyst concentration phase temperature was 720° C. The catalyst used in this experiment was a commercially-available catalyst HARMOREX (CCIC). Prior to feeding the catalyst into the device, the catalyst was subjected to steaming at 810° C. for 6 hours with 100% steam in order to bring the catalyst into a pseudo-equilibrium state. The feed oil was a hydrodesulfurized Arabian Light vacuum gas oil. With regard to the hydrotreating conditions of the vacuum gas oil, the hydrogen partial pressure was 150 kg/cm²G, the reaction temperature was 400° C, and the LHSV was 1. The vacuum gas oil, hydrotreated under the above conditions, had a specific gravity at 15° C. of 0.828 g/cm³, an aniline point of 118° C., a hydrogen content of 14.1 wt %, a nitrogen content of 0.003 wt %, and a sulfur content of 0.02 wt %. With regard to the distillation properties of the hydrotreated vacuum gas oil, the 10 vol % point was 387° C., 50 vol % point was 445° C., and 90 vol % point was 512° C. The results are shown in Table 1.

Comparative Example 1

Catalytic cracking of a heavy oil was carried out by the use of the same device and the same catalyst as Example 1. With regard to operation conditions, the reaction zone outlet temperature was 600° C., the reaction pressure was 1.0 kg/cm[0041] ²G, the catalyst/oil ratio was 14.9 wt/wt, the contact time was 0.4 seconds, and the regenerating zone catalyst concentration phase temperature was 720° C. The feed oil was a hydrodesulfurized Arabian Light vacuum gas oil. With regard to the hydrotreating conditions of the vacuum gas oil, the hydrogen partial pressure was 65 kg/cm²G, the reaction temperature was 400° C., and the LHSV was 2. The vacuum gas oil hydrotreated under the above conditions has a specific gravity at 15° C. of 0.897 g/cm³, an aniline point of 77.9° C., a hydrogen content of 12.6 wt %, a nitrogen content of 0.04 wt %, and a sulfur content of 0.13 wt %. With regard to the distillation properties of the hydrotreated vacuum gas oil, the 10 vol % point was 384° C., 50 vol % point was 462° C., and 90 vol % point was 556° C. The results are shown in Table 1.

TABLE 1

Example 1 Comparative Example 1

Conversion wt % 95.6 86.3

Yield wt %

Dry gas 4.1 3.8

Propylene 18.2 11.3

Butenes 22.5 15.0

Gasoline 42.5 48.7

LCO+ 4.4 13.7

Coke 1.0 2.7
It is apparent, from the above-mentioned results, that when a vacuum gas oil hydrotreated under high hydrogen partial pressure is used as a feed oil of the catalytic cracking apparatus operated at high reaction temperature and with short contact time, the conversion of the vacuum gas oil will be high, the yield of light olefins will be high, and the yield of coke will be high. [0042]
According to the process of this invention for fluid catalytic cracking of a heavy-fraction oil, the yield of light olefins can be heightened while the yield of coke can be lessened.[0043]

Claims

What is claimed is:

1. A process for the fluid catalytic cracking of a heavy-fraction oil, which comprises the step of contacting a vacuum gas oil, which has been treated with hydrogen under the hydrogen partial pressure higher than 80 kg/cm²G, with a catalyst in a fluid catalytic cracking apparatus having a regeneration zone, a reaction zone, a separation zone and a stripping zone, and under the conditions that the reaction zone outlet temperature is in the range of 550 to 630° C. and the contact time of hydrocarbons in the reaction zone is in the range of 0.01 to 1.0 sec.

2. The process according to claim 1 wherein the reaction zone outlet temperature is in a range of 550 to 630° C.

3. The process according to claim 1 wherein the contact time of hydrocarbons in the reaction zone is in a range of 0.05 to 0.8 seconds.

4. The process according to claim 1 wherein the contact time of hydrocarbons in the reaction zone is in a range of 0.1 to 0.6 seconds.

5. The process according to claim 1 wherein the fluid catalytic cracking apparatus is operated at a reaction pressure of 1 to 3 kg/cm²G.

6. The process according to claim 1 wherein the fluid catalytic cracking apparatus is operated with a catalyst/oil ratio of 8 to 40 wt/wt.

7. The process according to claim 1 wherein the fluid catalytic cracking apparatus is operated with a catalyst/oil ratio of 12 to 30 wt/wt.

8. The process according to claim 1 wherein the reaction zone is a down flow-type reaction zone.

9. The process according to claim 1 wherein the heavy-fraction oil is a vacuum gas oil which has been treated with hydrogen under hydrogen partial pressure higher than 145 kg/cm²G.

10. The process according to claim 8 wherein the vacuum gas oil has been treated with hydrogen at a reaction temperature of 350 to 430° C., at an LHSV of 0.5 to 4.0 and at a H₂/Oil of 1,000 to 4,000 scf/bbl.

11. The process according to claim 1 wherein the vacuum gas oil has a hydrogen content more than 13 wt %, a nitrogen content less than 0.01 wt %, a sulfur content less than 0.1 wt %, a specific gravity at 15° C. of 0.80 to 0.85 g/cm³, an aniline point of 100 to 130° C., a 10 vol % distillation point of 250 to 470° C., a 50 vol % distillation point of 300 to 520° C. and a 90 vol % distillation point of 340 to 560° C.

12. The process according to claim 10 wherein the hydrogen content is more than 13.5 wt %, the nitrogen content is less than 0.005 wt %, the 10 vol % distillation point is in the range of 340 to 420° C., the 50 vol % distillation point is in the range of 400 to 500° C., and the 90 vol % distillation point is in the range of 470 to 550° C.