HEAVY OIL CRACKING METHOD

Abstract

A method for cracking heavy oil is disclosed. The method uses a first heating stage, a second heating stage, a first cracking stage and a second cracking stage to produce cracked distillates from the residual heavy oil.

Description

CROSS REFERENCE TO RELATED APPLICATIONS

This application claims priority from U.S. Provisional Application Ser. No. 61/223,885 filed Jul. 8, 2009.

BACKGROUND OF THE INVENTION

The invention is directed to a method for cracking heavy oil. More particularly, the invention is directed to a method for cracking heavy oil using staging to convert the heavy oil feed to a cracked pitch product.

A recent review article [Hulet (2005)] examined the key features and configurations of short residence time cracking processes developed over the past 25 years. This work succinctly summarized the promise, key features, and challenges of short residence time processes:

“There is a strong economic incentive for considering short residence time cracking processes. Not only do such processes increase the yields of the more valuable liquid and gaseous products, but more compact designs would also decrease capital costs. Careful control of the vapor residence times appears to be crucial in order to prevent secondary cracking and yet allow for maximum cracking of the feedstock. Rapid and thorough mixing of the feedstock with the heat source, not just creating a uniform dispersion, is also a key design aspect to consider. Finally, rapid and complete separation must also be carefully considered; again, to help control product residence time and avoid secondary cracking but also from a heat balance point of view.”

These insights suggest short residence time cracking processes should rapidly heat the feed, rapidly separate products to control residence time of all products, rapidly cool the reaction products to avoid secondary cracking, and make productive use of the thermal energy. The most successful short residence time processes meet all these criteria. However, none of these processes fully meet all these criteria while treating asphaltic residual oils. Fully meeting these criteria with asphaltic residual oils is the goal for more effective deasphalting, thermal cracking, and hydrocracking processes.

Fluid catalytic cracking (FCC) is undoubtedly the most common and commercially successful short residence time cracking process. Typically, the FCC process intimately contacts a gas oil boiling range hydrocarbon feedstock with hot catalyst particles in an entrained flow, short-residence, riser reactor to produce more valuable cracked products, particularly gasoline and olefins, and less desirable dry gas and coke by-products. The FCC process developers have used improved feed nozzle designs to increase feed heating rate. The FCC feed nozzles improve the uniformity of the initial contact between the carbonaceous feed and the hot regenerated catalyst, which increases the feed heating rate and decreases the yield of the undesirable dry gas and coke FCC produces. For example, U.S. Pat. No. 6,387,247, summarizes a long standing effort to use feed injection nozzle improvements to increase the feedstock heating rate and improve the overall FCC reactor performance.

FCC process developers have also identified approaches to control the residence time. For example, U.S. Pat. No. 6,979,360, teaches methods for conducting short contact time hydrocarbon conversions in a FCC reactor with the rapid inertial separation of gas and solid FCC reactor products. U.S. Pat. No. 6,616,900 extends this concept by using a staged FCC riser reactor with interstage product removal. U.S. Pat. No. 5,762,882, teaches methods to remove reaction products from the spent catalyst via vaporization. With distillate feeds, the by-product coke production is roughly in balance with FCC process heat requirement. The higher coke yield associated with more asphaltic residual FCC feeds has a large adverse effect on the process performance. U.S. Pat. No. 4,415,438 teaches the use of a thermally stable catalyst and high catalyst regeneration temperature to increase the heavy oil feed heating rate and decrease coke yield. U.S. Pat. No. 5,271,826 achieves a similar result by increasing the regenerated catalyst to feed ratio to achieve an elevated riser initial temperature and then adding a quench liquid to temperature the riser temperature. Canadian Patent No. 2,369,288, teaches thermal cracking of residual oil feed with inert solids in an FCC reactor-type short contact time reactor to eliminate catalyst deactivation problems, but also results in an inferior product yield distribution, including coke production in excess of the process heat requirement. US Patent Application Publication No. 2006/0042999 teaches deasphalting of the FCC heavy oil feed to decrease the catalyst deactivation rate due to metals and coke precursors in the feed. U.S. Pat. No. 6,171,471 teaches the combination of mild hydrocracking and deasphalting of the residual oil FCC feed to decrease the catalyst deactivation rate due to metals and coke precursors. Despite these efforts to decrease the FCC process coke yield with residual asphaltic feeds, the coke yield far exceeds the amount required to preheat the regenerated catalyst or inert solids.

As a result, similar principles were used to develop successful short contact time fluid coking processes that maximize the conversion of asphaltic residual oil feeds to distillates. The fluid coking process typically comprises partial combustion of coke particles, rapid heating of the residual oil feed by intimate contact with a fluidized bed of hot coke particles, rapid separation of entrained by-product coke from the vapor product using cyclones, and a quench system to rapidly cool the vapor product to minimize secondary thermal cracking. U.S. Pat. No. 2,881,130 teaches atomization and distribution of the residual oil feed to prevent bogging of the fluidized coke bed and to increase the residual oil feed heating rate. The fluid coking process developers have also identified methods to control residence time. For example, U.S. Pat. No. 4,816,136 teaches sequentially contacting the feed with higher temperature coke particles in a riser reactor and then in a fluidized coke bed that operates at a lower temperature to increase the distillate yield. U.S. Pat. No. 5,658,455 describes a method for a short vapor residence time reactor to minimize secondary thermal cracking reactions. U.S. Pat. No. 4,497,705 teaches methods to solvent refine the recycle heavy oil to selectively remove less carbonaceous species to decrease undesirable secondary cracking of these valuable products. U.S. Pat. No. 4,587,010 teaches methods to strip valuable products from the coke prior to regeneration via partial oxidation. The fluid coking has many common features with the FCC process and has several advantages for treating carbonaceous residual oil feeds. The fluid coking process eliminates the rapid catalyst deactivation problem and need to burn very large quantities of coke that are associated with the FCC process treating asphaltic residual oil feeds. Since both the fluid coking and FCC processes require that all liquid products are produced by vaporization, neither process can operate with a short residence time for unconverted asphaltic residual oils.

U.S. Pat. No. 3,393,133 teaches high temperature and short residence time distillation processes to maximize distillate yield with minimum degradation of the residual oil due to thermal cracking reactions. However, solvent extraction [Altgelt (1994)] is the preferred method to produce residual oil fractions with much higher equivalent normal boiling points and essentially no thermal cracking degradation. Successful solvent refining processes have been developed, e.g. U.S. Pat. No. 4,810,367, to continuously produce deasphalted oil, resin, and asphaltene streams from an asphaltic residual oil feed. These processes require a large number, but simple and reliable, unit operations to contact and separate the residual oil from the solvents. In addition, the solvent and residual oil separation steps have a significant steam heat requirement. As a result, this process technique is particularly useful in petroleum refineries, where the required solvents, steam, and maintenance infrastructure are readily available. U.S. Pat. No. 6,357,526 teaches a solvent extraction field upgrader method to produce a deasphalted oil synthetic crude product and an asphaltic fuel to produce steam for bitumen extraction. For this remote application, a flash deasphalting process has many potentially desirable features. A very high temperature, very short contact time flash unit operation could produce the deasphalted oil and asphaltic streams in a compact and single unit operation without the need for a solvent. The thermal energy input could be used to separate the deasphalted oil and asphaltene products, produce steam for bitumen extraction, and produce a hot asphaltene stream that can be burned without the need for reheating or pelletization. Unfortunately, the earlier processes do not provide any method that can heat the bitumen feed, separate the deasphalted oil and asphaltene products, and cool the separated deasphalted oil and asphaltene products sufficiently rapidly to avoid excessive thermal cracking and degradation of the deasphalted oil product.

Resid hydrocracking is the most well established method to convert asphaltic materials to less carbonaceous materials and to reduce the metals and coke precursor concentration in the unconverted asphaltene species. U.S. Pat. No. 2,987,465 first introduced the ebullated bed hydrocracking reactor concept. An ebullated hydrogenation reactor utilizes up-flow of the carbonaceous asphaltic residual oil and hydrogen feeds to contact an expanded bed of particulate hydrotreating catalyst and or entrained colloidal hydrotreating catalyst. The expanded bed hydrotreating catalyst bed is much less susceptible to plugging than the previous fixed catalyst bed designs. U.S. Pat. No. 5,164,075 teaches methods to produce colloidal heavy oil catalysts that are particularly effective for hydrogenating asphaltic species. U.S. Pat. No. 6,511,937 teaches methods to recover and recycle colloidal heavy oil hydrocracking catalysts. All these resid hydrocracking process simultaneously hydrogenate and thermally crack the residual oil. U.S. Pat. No. 4,427,535 identifies a fundamental limitation with this approach. The thermal cracking reactions have higher activation energies than the hydrogenation reactions. Hydrogenation reactions retard polymerization reactions that produce coke precursors and ultimately coke. Therefore, a resid hydrocracker must be operated at a much lower temperature than a FCC or fluid coking unit in order to maintain operability.

US Patent Application Publication No. 2005241993 teaches the use of a colloidal molybdenum sulfide catalyst to increase the hydrogenation rate, particularly the rate of hydrogenation of asphaltic species. This innovation increases hydrogenation rate of asphaltic species and the maximum operable temperature, but does not alter the basic nature of this hydrocracker operating temperature limitation. Hydrogen donor diluent cracking processes substantially eliminates this temperature limitation by hydrogenating a naphthenic distillate at moderate temperatures to produce a hydrogen donor solvent and then thermal cracking the residual oil in the presence of the donor solvent to substantially reduce the rate of coke precursor formation. U.S. Pat. No. 4,698,147 teaches simultaneously increasing the hydrogen donor diluent cracking temperature and decreasing the contact time to monotonically increase the maximum resid conversion to distillates. U.S. Pat. No. 4,002,556 teaches that high temperature and short contact time hydrogen donor diluent cracking also substantially reduces the hydrogen consumption required to maintain operability. These high temperature and short contact time benefits seem to be only limited by the practical limits on the feed heating rate and the product cooling rate. The hydrogen donor diluent process requires sufficient pressure to maintain a liquid phase hydrogen donor solvent at the elevated thermal cracking temperatures.

The invention is able to address the concerns of these earlier processes by using a staging process to convert heavy oil into a cracked distillate.

SUMMARY OF THE INVENTION

The invention provides for advantages to the process of converting residual heavy oil. The invention improves control over the reactor composition. The staging process provides for prompt removal of cracked distillate product to inhibit secondary cracking and increases the solubility of coke precursor in the liquid phase. Thermal efficiencies are also realized by the invention.

The invention further provides for a method for cracking heavy oil comprising feeding residual heavy oil first to a heating stage comprising an atomizer, a contactor and an inertial vapor-liquid separator, and then to at least one stage selected from the group consisting of temperature maintenance stage comprising a reactor and an inertial vapor-liquid separator, and a second heating stage comprises an atomizer, a contactor and an inertial vapor-liquid separator

The invention provides for a method of cracking residual heavy oil comprising at least one heating stage and at least one or more of a stage selected from the group consisting of a heating stage and a temperature maintenance stage.

The heating stage is accomplished by three unit operations comprising an atomizer, a contactor and an inertial vapor-liquid separator.

The temperature maintenance stage is accomplished by two unit operations comprising a reactor and an inertial vapor-liquid separator.

The method of cracking residual heavy oil can be accomplished by any combination of the two stages can be employed so two heating stages may be employed or two temperature maintenance stages or one or more of each stage. Thus numerous combinations of heating stages and temperature maintenance stages can be employed as long as there is more than one stage used in the cracking process.

The heating stage starts with a feed of a heating gas formed from the reaction products of an oxygen-containing compound and a fuel. This reaction occurs in a heating pressurized combustor. The heating gas is fed to the atomizer along with a residual heavy oil feed.

The now atomized residual heavy oil is fed to a contactor where the residual heavy oil is maintained for sufficient residence time to achieve the desired conversion of residual oil to distillates via thermal cracking reactions.

The products from the contactor are then fed to the inertial vapor-liquid separator where the products are separated into vapor and liquid portions.

The vaporized gas and liquid distillate products are removed from the inertial vapor-liquid separator and the remaining residual oil liquid is fed to the second stage of the method, the temperature maintenance stage.

The first step of the temperature maintenance stage is the reactor step where the reactor receives the liquid product from the inertial vapor-liquid separator of the heating stage. The reactor will provide sufficient residence time to achieve the desired cracking.

The cracked reaction products are then fed to an inertial vapor-liquid separator where they will mix with the reaction products from a cracking pressurized combustor. The liquid portion and vapor portion are separated and the gas and cracked distillate product recovered.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a schematic view of a pressurized combustor.

FIG. 2A is a schematic view of a conical atomizer.

FIG. 2B is a schematic view of a cylindrical atomizer.

FIG. 3 is a schematic view of a vapor-liquid contactor.

FIG. 4 is a schematic view of an inertial vapor-liquid separator.

FIG. 5 is a schematic view of a two stage heating and thermal cracking process according to the invention.

DETAILED DESCRIPTION OF THE INVENTION

The invention will now be described in more detail and with reference to the drawing figures.

The pressurized combustor on FIG. 1 mixes a fuel 1 with an oxidant 2 in a burner 3. The fuel may be any hydrocarbon and/or hydrogen. The fuel is preferably a H₂ and CO synthesis gas that is produced by conventional gasification of the pitch by-product (see Stream 40 on FIG. 5). The oxidant 2 may be any mixture of air, O₂, and steam. The burner 3 utilizes conventional ignition and flame monitoring methods to maintain a stable flame 4. The combustion reactions are substantially completed within a pressurized shell 5. The pressurized shell 5 may be advantageously fitted with internal insulation 6 to decrease heat loss and the pressurized combustor shell 5 temperature. The pressurized combustor product gas 7 would typically be in the 5 to 20 bar pressure range, 1400-1800° C. temperature range, and 0 to −5% excess O₂.

Typical heavy oil atomizers are illustrated on FIGS. 2A and 2B. FIG. 2A illustrates a conical convergent-divergent nozzle atomizer and FIG. 2B illustrates a cylindrical convergent-divergent nozzle atomizer. The pressurized combustor product gas 7 is fed to the nozzle motive plenum 8. Similar and well established methods can be used to design both types of nozzles in terms of the motive gas plenum 8 temperature, pressure, and composition, the nozzle throat 9 area, the expansion angle 10, and nozzle 11 discharge area.

In the case of the conical nozzle (see FIG. 2A), the atomization nozzle heavy oil feed 12 is directed by the heavy oil conduit 13 to the conical nozzle mixing chamber 14. The heavy oil atomization effectiveness is a complicated function of the conical gas jet 15 kinetic energy and temperature, conical gas jet intersection angle 16, and conical nozzle mixing chamber 14 volume and shape. Generally, the conical atomization nozzle efficiency increases with increasing the conical gas jet 15 kinetic energy and temperature, conical gas jet intersection angle 16, and decreasing conical nozzle mixing chamber 14 volume. The atomized heavy oil product 17 from the conical nozzle is fed to the vapor-liquid contactor or reactor 18 (see FIG. 3).

The vapor-liquid contactor or reactor 18 is illustrated on FIG. 3. This vessel is considered a vapor-liquid contactor if its primary function is to provide the atomized heavy oil sufficient residence time to approach the temperature of the gas phase that is substantially below design thermal cracking temperature, typically between 650° C. and 850° C. This vessel is considered to be a reactor if the heavy oil temperature is near the design thermal cracking temperature. The vapor-liquid contactor or reactor 18 is typically a cylindrical pressure vessel with a feed inlet nozzle 19 and a product outlet nozzle 20. The vapor-liquid contactor or reactor 18 may be advantageously equipped with internal insulation 21 to decrease the heat loss and the vapor-liquid contactor or reactor 18 shell temperature. The diameter of the vapor-liquid contactor or reactor 18 diameter is set to provide highly turbulent flow with a reasonable pressure drop. The vapor-liquid contactor or reactor 18 length is set to achieve the desired gas less atomized heavy oil temperature difference when operating in the vapor-liquid contactor operating mode or the desired thermal cracking extent of reaction prior to separating the cracked vapor product from the unconverted heavy oil when operating in the reactor operating mode.

The preferred inertial vapor-liquid separator is a cyclone. FIG. 4 illustrates the key features of the cyclone inertial vapor-liquid separator. The general shape of the cyclone inertial separator 22 is a truncated cone with the smaller diameter at the feed end 23 and the larger diameter at the product end 24. The product end 24 diameter would generally be between 1.1 and 1.4 times the feed end 23 diameter to ensure a low liquid inventory and residence time in the vapor-liquid separator. Internal insulation (not shown on FIG. 4) may be advantageously used to decrease inertial separator heat loss and wall temperature. A rectangular conduit 25 is preferably used to shape the inertial vapor-liquid separator feed 26. The inertial vapor-liquid separator is advantageously fitted with a weir 27 to minimize short-circuiting of liquid droplets in the feed 26 to the vapor outlet 28. A distillate quench stream 29 may be advantageously used to decrease the vapor product 30 temperature to between 300° and 400° C. to control the rate of thermal cracking reactions in the vapor product 30.

An atomizer is advantageously used to increase the interfacial area of the hydrocarbon quench liquid 29. The hydrocarbon quench liquid preferably has a normal boiling point less than 350° C. The liquid product conduit 32 is used to remove the liquid product 33 from the inertial vapor-liquid separator. A liquid product conditioning stream 34 may be advantageously introduced via the liquid product conditioning stream conduit 35 to achieve the desired liquid product 33 properties. For example, one could advantageously use steam as the liquid product conditioning stream 34 to minimize the quantity of the vapor in the feed stream 26 that exits in the liquid product stream 33 and cools the liquid product stream 33 to control the thermal cracking rate. Alternatively, one could advantageously use the pressurized combustor product gas 7 on FIG. 1 to control the quantity of the vapor in the feed stream 26 that exits in the liquid product stream 33 and heat the stream 33 to compensate for heat losses and the endothermic thermal cracking heat of reaction. Typically, the inertial separator feed 26, the liquid product conditioning stream 34 would all have either a clockwise or counterclockwise flow. One can also advantageously improve liquid drainage and decrease the vapor-liquid separator volume by using a product end insert 36. The relative dimensions in FIG. 4 are only intended to provide rough guidance for the inertial vapor-liquid separator. The vapor residence time in the inertial vapor-liquid separator would typically be between 1 and 5 milliseconds.

This process description will describe the options to use these unit operations for a staged thermal cracking with the aid of FIG. 5.

The stage cracking method (See FIG. 5) seeks to maximize the conversion of the residual heavy oil feed 37 to substantially uncracked distillate products 38 and cracked distillate products 39 and minimize the pitch product yield. The staged heavy oil cracking system on FIG. 5 utilizes two heating stages and two thermal cracking stages. This staged method consists of at least one preheating and one thermal cracking stage. A more typical arrangement would consist of two heating and two thermal cracking stages, as shown on FIG. 5. Up to four heating and cracking stages can be advantageously used. A heating stage would typical consist of a conical atomizer (see FIG. 2A), a contactor vessel (see FIG. 3), and a cyclone inertial vapor-liquid separator (See FIG. 4). The residual heavy oil feed 37 and 1^st heating stage atomizing and heating gas 42 are introduced into the 1^st heating stage atomizer 41.

The asphaltic residual oil feed 37 typically contains more than 25 weight percent of species with normal boiling points greater than 975° F. (524° C.), more preferably greater than 50 weight percent, and most preferably greater than 75 weight percent. The residual oil feed typically has micro carbon and heptane insoluble contents between 5 and 40 weight percent. The more carbonaceous feeds typically exhibit higher micro carbon and heptane insoluble values. Typical sources for residual oil feed include petroleum atmospheric or vacuum residual oil, oil sands, bitumen, tar sand oils, coal tar, pyrolysis tars, or shale oils.

The residual oil feed 37 may also be partially hydrogenated prior entering atomization-heating nozzle 41. The optional residual oil feed hydrogenation step typically has a hydrogen consumption between 100 and 1500 standard cubic feet per petroleum barrel (SCF/bbl) or between 95 and 285 gram moles per cubic meter, more preferably between 150 and 1000 SCF/bbl, most preferably between 200 and 700 SCF/bbl. The optional hydrogenation step is preferably performed in an ebullated bed hydrotreating reactor with conventional cobalt-molybdate or nickel-molybdate on activated alumina supported hydrotreating catalyst and/or colloidal hydrogenation catalyst system, preferably a molybdenum sulfide colloidal catalyst.

The residual oil feed 37 is preheated in a conventional fired heater (not shown on FIG. 5) to the maximum temperature consistent with modest thermal cracking rates, typically in the 350 to 450° C. temperature range.

The 1^st heating stage atomizing and heating gas 42 is typically produced in the heating stage pressurized combustor using air oxidant 44 with optional steam additions to prevent sooting and a convert fuel 45. Alternatively, one could use an O₂-steam oxidant. The O₂ flow rate in the oxidant 44 feed to the pressurized combustor 43 is typically between 90% and 100% of the theoretical flow rate required to convert the hydrogen and/or carbon in the fuel feed 45 to CO₂ and H₂O. The 1^st heating stage atomizing and heating gas 42 typically is in the temperature range of 1300-1800° C. and the pressure range of 5 to 20 bar. The heating stage pressurized combustor fuel 45 may be any hydrocarbon fuel or hydrogen. The preferred fuel is a H₂ and CO synthesis gas produced by gasification of the pitch product 40.

The 1^st heating stage atomizer 41 (see FIG. 2A) uses the 1^st heating stage atomizing and heating gas 42 to atomize the residual heavy oil feed 37 and 1^st heating stage contactor 46 (See FIG. 3) provides sufficient residence so that the oil droplets approach the gas temperature, typically within 10-20° C., which would generally be accomplished within a 1-10 millisecond time frame.

The 1^st heating stage contactor product 47 is promptly fed to the 1^st heating stage inertial separator 48 (see FIG. 4). The 1^st heating stage inertial separator 48 rapidly (typically less than 5 milliseconds) separates the vapor and liquid portions of the 1^st heating stage contactor product 47. The 1^st heating stage inertial separator quench 49 promptly reduces the 1^st heating stage inertial separator distillate product 50 to about 350° C. to control the extent of thermal cracking of the distillate product 50.

The operation of the 2^nd heating stage atomizer 52, contactor 54, and inertial separator 56 are similar to the operation of the corresponding unit operations in the 1^st heating stage with some minor exceptions. First, the ratio of the flow rate of the atomizing and heating gas for the 1^st stage 42 is typically substantially higher than the flow rate for the 2^nd stage 53. Generally, these flow rates are adjusted such that the temperature of the 2^nd heating stage inertial separator liquid product 59 has reached the desired thermal cracking temperature, typically in the 650 to 850° C. temperature range, and the vaporized distillate flow rates in streams 50 and 58 are roughly equivalent. In addition, the residence time in the 2^nd heating stage contactor 54 is less than the 1^st heating stage contactor in order to control the extent of thermal racking. As a result, the temperature difference of the vapor and liquid portion of the 2^nd heating stage contactor product 55 is greater than the 1^st heating stage contactor product 47.

1^st cracking stage reheat and purge gas 60 is typically produced in the cracking stage pressurized combustor 61 using a steam-O₂ oxidant 62 and a convent fuel 63. The O₂ flow rate in the oxidant 62 feed to the cracking pressurized combustor 61 is typically between 90% and 100% of the theoretical flow rate required to convert the hydrogen and/or carbon in the fuel feed 63 to CO₂ and H₂O. The 1^st cracking stage reheat and purge gas 60 typically is in the temperature range of 1000-1600° C. and the pressure range of 2-10 bar. The lower temperature and pressure of the 1^st cracking stage reheat and purge gas 60 relative to 1^st heating stage atomizing and heating gas 42 results from the fact that this stream must only provide sufficient thermal energy to compensate for modest heat losses and heat of cracking in order to maintain the desired thermal cracking temperature. This lower heating requirement also eliminates the need for an atomizer. The cracking stage pressurized combustor fuel 63 may be any hydrocarbon fuel or hydrogen. The preferred fuel is a H₂ and CO synthesis gas produced by gasification of the pitch product 40. A water quench stream 64 may be advantageously used to obtain the desired 1^st cracking stage reheat and purge gas 60 temperature.

The 1^st cracking stage reheat and purge gas 60 in fed to the terminal reheating inertial separator 56 as shown on FIG. 4 and described above in order to minimum the nitrogen and uncrack distillate content of the 2^nd heating stage inertial separator liquid product 59. A minority, typically less than 20%, of the 1^st cracking stage reheat and purge gas exits 2^nd heating stage inertial separator 56 via stream 58. The 2^nd heating stage inertial separator liquid product 59 is promptly fed to the 1^st cracking stage reactor 65. The diameter of the 1^st cracking stage reactor 65 is adjusted to provide a highly turbulent flow regime in the 1^st cracking stage reactor 65 to minimize liquid hold-up on the wall and to achieve reasonable heat transfer from the gas to the liquid phase. The length of the 1^st cracking stage reactor 65 is adjusted to provide sufficient residence time to achieve the desired conversion. The 1^st cracking stage reactor product 66 is feed to the 1^st cracking stage inertial separator 67, which is operated in the same manner at the 2^nd heating stage inertial separator 56. Likewise, the 2^nd cracking stage is operated in the same manner as the first stage except a pitch product 78 quench, usually water, can be advantageously used to reduce the pitch product 40 temperature to about 350° C. and minimize further degradation.

As noted in FIG. 5, the second stage cracking reheat and purge gas enters cyclone 67 through line 71. Gas and cracked pitch product will exit cyclone 67 through line 70 to reactor 72 where it will continue through line 73 to cyclone 74. The second cracking stage distillate quench is fed through line 75 to cyclone unit 74. This feed is accompanied by the feed from reactor 72 which enters cyclone 74 through line 73. Gas and cracked distillate product leave through line 76 to line 39 and exit the system. The gas and cracked pitch product will exit through line 77 where it is quenched by a pitch product quench entering through line 78. The now quenched gas and cracked pitch product will exit the system through line 40.

While this invention has been described with respect to particular embodiments thereof, it is apparent that numerous other forms and modifications of the invention will be obvious to those skilled in the art. The appended claims in this invention generally should be construed to cover all such obvious forms and modifications which are within the true spirit and scope of the present invention.

Claims

1. A method for cracking heavy oil comprising at least one heating stage and at least one stage selected from the group consisting of a second heating stage and a temperature maintenance stage.

2. The method as claimed in claim 1 wherein said at least one stage is a second heating stage.

3. The method as claimed in claim 1 wherein said at least one stage is a temperature maintenance stage.

4. The method as claimed in claim 1 wherein up to four of each one at least one heating stage and at least one stage selected from the group consisting of a second heating stage and a temperature maintenance stage are present.

5. The method as claimed in claim 1 wherein said heating stage occurs in a heating pressurized combustor.

6. The method as claimed in claim 1 wherein an oxidation reaction occurs in said heating stage.

7. The method as claimed in claim 6 wherein heating gas formed in said heating pressurized combustor is fed to an atomizer along with a residual heavy oil feed.

8. The method as claimed in claim 7 wherein said residual heavy oil feed contains at least 25% of species having a boiling point great than 524° C.

9. The method as claimed in claim 8 wherein said residual heavy oil is partially hydrogenated prior to atomization.

10. The method as claimed in claim 9 wherein said atomized residual heavy oil feed is fed to a contactor and maintained for a residence time sufficient to convert the atomized residual heavy oil to distillates.

11. The method as claimed in claim 10 wherein reactions occurring in said contactor are at a temperature between 650° and 850° C.

12. The method as claimed in claim 6 wherein reaction products from said contactor are fed to an inertial vapor-liquid separator to separate said reaction products into vapor and liquid portions.

13. The method as claimed in claim 12 wherein said inertial vapor-liquid separator is a cyclone.

14. The method as claimed in claim 13 wherein remaining residual oil liquid is fed to said temperature maintenance stage.

15. The method as claimed in claim 14 wherein said temperature maintenance stage is a reactor for cracking said residual oil liquid.

16. The method as claimed in claim 15 wherein said temperature maintenance stage is at a temperature of 1000° to 1600° C. and 2 to 10 bar.

17. The method as claimed in claim 15 wherein said cracked residual oil reaction products are fed to said inertial vapor-liquid separator containing the distillates.

18. The method as claimed in claim 10 wherein said cracked residual oil reaction products and distillates are separated into a liquid portion and a vapor portion.

19. The method as claimed in claim 1 wherein said heating stage comprises an atomizer, a contactor and an inertial vapor-liquid separator.

20. The method as claimed in claim 19 wherein said atomizer is selected from the group consisting of conical and cylindrical convergent-divergent nozzle atomizers.

21. The method as claimed in claim 1 wherein said temperature maintenance stage comprises a reactor and an inertial vapor-liquid separator.

22. The method as claimed in claim 1 wherein said second heating stage comprises an atomizer, a contactor and an inertial vapor-liquid separator.

23. A method for cracking heavy oil comprising feeding residual heavy oil first to a heating stage comprising an atomizer, a contactor and an inertial vapor-liquid separator, and then to at least one stage selected from the group consisting of temperature maintenance stage comprising a reactor and an inertial vapor-liquid separator, and a second heating stage comprises an atomizer, a contactor and an inertial vapor-liquid separator

24. The method as claimed in claim 23 wherein said at least one stage is a second heating stage.

25. The method as claimed in claim 23 wherein said at least one stage is a temperature maintenance stage.

26. The method as claimed in claim 23 wherein up to four of each one at least one heating stage and at least one stage selected from the group consisting of a second heating stage and a temperature maintenance stage are present.

27. The method as claimed in claim 23 wherein said heating stage occurs in a heating pressurized combustor.

28. The method as claimed in claim 23 wherein an oxidation reaction occurs in said heating stage.

29. The method as claimed in claim 28 wherein heating gas formed in said heating pressurized combustor is fed to an atomizer along with a residual heavy oil feed.

30. The method as claimed in claim 29 wherein said residual heavy oil feed contains at least 25% of species having a boiling point great than 524° C.

31. The method as claimed in claim 30 wherein said residual heavy oil is partially hydrogenated prior to atomization.

32. The method as claimed in claim 23 wherein said atomized residual heavy oil feed is fed to a contactor and maintained for a residence time sufficient to convert the atomized residual heavy oil to distillates.

33. The method as claimed in claim 32 wherein reactions occurring in said contactor are at a temperature between 650° and 850° C.

34. The method as claimed in claim 32 wherein reaction products from said contactor are fed to an inertial vapor-liquid separator to separate said reaction products into vapor and liquid portions.

35. The method as claimed in claim 23 wherein said inertial vapor-liquid separator is a cyclone.

36. The method as claimed in claim 29 wherein remaining residual oil liquid is fed to said temperature maintenance stage.

37. The method as claimed in claim 23 wherein said temperature maintenance stage is a reactor for cracking said residual oil liquid.

38. The method as claimed in claim 23 wherein said temperature maintenance stage is at a temperature of 1000° to 1600° C. and 2 to 10 bar.

39. The method as claimed in claim 23 wherein said cracked residual oil reaction products are fed to said inertial vapor-liquid separator containing the distillates.

40. The method as claimed in claim 32 wherein said cracked residual oil reaction products and distillates are separated into a liquid portion and a vapor portion.

41. The method as claimed in claim 23 wherein said atomizer is selected from the group consisting of conical and cylindrical convergent-divergent nozzle atomizers.