FLASH PROCESSING OF ASPHALTIC RESIDUAL OIL

Abstract

A method to upgrade virgin and partially hydrogenated asphaltic residual oils by utilizing hot, high velocity combustion gas jets to rapidly atomize and heat the residual oil, maintaining the reactant temperature required to achieve the desired residual oil conversion with the minimum practical residence time, rapidly separating vapor and liquid reactants, and rapidly cooling the vapor and liquid products. The minimum required temperature and practical residence time are used for the production of deasphalted oil and asphaltene products with minimum degradation due to thermal cracking. The maximum conversion of residual oil may be substantially increased by combining a portion of the heavy oil product with the residual oil feed and partially hydrogenating this mixture.

Description

BACKGROUND OF THE INVENTION

The general field of this invention is high temperature and short contact time processing of asphaltic residual oil. More specifically, this invention is a flash process method to remove asphaltic species from residual oils.

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. An asphaltic residual oil has a significant concentration of species with normal boiling points greater than 524° C. (typically greater than 25 wt %) and heptane insoluble specie concentrations greater than twenty-five weight percent of the species with normal boiling points greater than 524° C. 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.

None of these processes can achieve the potential advantages of a high temperature short contact process that produces a liquid asphaltic product per the methods of the present invention.

SUMMARY OF THE INVENTION

The present invention is a flash method to upgrade asphaltic residual oils by utilizing combustion gas jets to rapidly atomize and heat the residual oil feed, using an inertial device to rapidly separate vapor and liquid reactants, and a quench stream to rapidly cool the vapor light oil and the liquid heavy oil products.

The combustion jet produced by the expansion of combustion gases through a convergent-divergent nozzle, also known as de Laval-type nozzle. The reactant residence time above 425° C. is less than 400 milliseconds, preferably less than 100 milliseconds, most preferably less than 50 milliseconds. Since asphaltic carbonaceous materials, i.e. heptane insoluble species are most the problematic species for conventional refinery unit operations, the present invention focuses on removal of asphaltic materials. The asphaltic carbonaceous materials can be removed from a residual oil by heating the reactants to the minimum temperature required to vaporize the heptane soluble species with the minimum residence time required for vaporization to produce deasphalted oil and asphaltic heavy oil products with minimum degradation due to thermal cracking. The asphaltic heavy oil product yield can be substantially reduced, with minimum gas and light distillate yield, by increasing the reactant temperature, while maintaining the minimum practical residence time above 425° C. The asphaltic heavy oil yield may be further decreased by partial hydrogenation of the residual oil feed and operating with a higher reactant temperature, while maintaining the minimum practical residence time above 425° C.

In one embodiment of the present invention, there is disclosed a method for upgrading asphaltic residual oil comprising:

• a) producing a combustion gas jet by expansion of a combustion gas through a convergent-divergent nozzle;
• b) contacting intimately the combustion gas jet with the asphaltic residual oil feed;
• c) separating the vapor phase and the liquid phase to separate the products formed by contacting the asphaltic residual oil with the combustion gas jet; and
• d) cooling the vapor phase and the liquid phase products to form light and heavy oil products, respectively.

The methods of the present invention operate when asphaltic residual oil is pretreated to a temperature between about 100° and 425° C. prior to feeding to said gas jet.

The vapor phase and the liquid phase products are cooled to temperatures of less than 400° C.

The asphaltic residual oil generally has an apparent viscosity less than 1000 centipoises (cP).

The method further provides for pretreatment comprising partial hydrogenation of a carbonaceous residual oil. This hydrogenation can be performed in an ebullated or fixed bed hydrotreating reactor, in the presence of a catalyst which is selected from the group consisting of cobalt-molybdate and nickel-molybdate on activated alumina support. The catalyst may also be a colloidal hydrotreating catalyst such as a colloidal molybdenum sulfide type catalyst.

Hydrogen is added to the residual oil feed at about 100 to about 1500 standard cubic feet per barrel of residual oil feed with about 150 to about 1000 standard cubic feed of hydrogen per barrel of residual oil feed preferred.

The combustion gas is produced by reacting an oxidant with a fuel wherein the fuel is selected from the group consisting of carbon monoxide, hydrogen, gaseous hydrocarbons and mixtures thereof and the oxidant is selected from the group consisting of air, oxygen-enriched air and substantially pure oxygen. Steam may also be used to control combustion temperature, discourage coke formation, and simplify recovery product recovery system.

The combustion gas is typically at a temperature of about 1250 to about 2000° C. and a pressure of about between 2 and 40 bar.

The convergent-divergent nozzle may be axisymmetric, and the carbonaceous residual oil can contact the gas jet from the periphery of the gas jet. The combustion gas jet may also be produced by a conical convergent-divergent nozzle. The residual oil may contact the gas jet along the axis of the conical convergent-divergent nozzle.

In the methods of the present invention, the inertial device may be a cyclone. The reactant residence time above 425° C. is less than 400 milliseconds. A liquid distillate quench may be employed to rapidly cool the vapor phase and liquid phase reactants. In some instance, the liquid distillate quench is atomized. A purge gas may be employed to minimize liquid distillate quench loss in the heavy oil product and this purge gas is preferably steam.

BRIEF DESCRIPTION OF DRAWINGS

FIG. 1 is a schematic of a flash processing of asphaltic residual oils according to the methods of the present invention.

FIG. 2 is a diagram of the atomization heating nozzle.

FIG. 3 is a diagram of the inertial vapor-liquid separation device.

DETAILED DESCRIPTION OF THE INVENTION

The flash processing method for treating residual asphaltic oils is described with the aid of the figures. In FIG. 1, the asphaltic residual oil feed 1 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 Ramsbottom carbon and heptane insoluble contents between 5 and 40 weight percent. The more carbonaceous feeds typically exhibit higher Ramsbottom 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 1 may also be partially hydrogenated prior entering atomization-heating nozzle 2. 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 or fixed 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 hydrotreated or untreated residual oil feed 1 is typically preheated to a temperature between 100 and 425° C. prior to feeding to the atomization-heating nozzle 2, more preferably between 200 and 400° C., most preferably between 300 and 375° C. At the feed temperature, the residual oil feed has an apparent viscosity preferably less 1000 centipoise (cP) or 1 Pascal-seconds (Pa·s), more preferably less than 100 cP, and most preferably less the 10 cP. The atomization-heating nozzle 2 uses a combustion gas jet 3 to rapidly atomize and heat the residual oil feed 1 to produce the reactant stream 6. The combustion gas jet 3 is formed by expansion of the combustion gas 4 through a convergent-divergent nozzle 5. Standard methods (Green-1999) can be used to design the convergent-divergent nozzle 5 achieve the desired combustion gas jet 3 kinetic and thermal energy content. The combustion gas 4 is produced by combustion of a fuel 7 with an oxidant 8 in an insulated 9 combustion chamber 10. The fuel 7 may be carbon monoxide, hydrogen, solid carbonaceous fuels, liquid hydrocarbons, gaseous hydrocarbons, or their mixtures. Gaseous and liquid fuels are preferred. Coal, coke, petroleum coke are examples of solid carbonaceous fuels. The oxidant may be air, O₂-enriched air, or substantially pure oxygen with a steam diluent. Substantially pure oxygen has an oxygen concentration preferably greater than 0.7 molar fraction (i.e. argon and nitrogen molar fractions are less than 0.3), more preferably greater than 0.85 molar fraction, most preferably greater than 0.9 molar fraction. Air is the preferred oxidant 8 when the process objective is to vaporize a substantial portion of the reactant stream 6 with minimum thermal cracking degradation. Substantially pure oxygen, with a steam diluent, is the preferred oxidant 8 when the process objective is to both vaporize and thermally crack the reactant stream 6.

An appropriate conventional burner 11 is used to mix and ignite the fuel 7 and oxidant 8. The oxidant 8 flow rate is adjusted such the flow rate is between 0.90 and 1.05 times the value theoretically required to convert the fuel 7 to CO₂, H₂O, H₂S, and N₂, more preferably between 0.95 and 1. The fuel 7 and oxidant 8 properties are adjusted to achieve a combustion gas 4 temperature between 1000 and 2500° C., more preferably between 1250 and 2000° C., and most preferably between 1500 and 1750° C. The combustion gas 4 pressure is preferably between 2 and 20 times the pressure of the reactant stream 6, more preferably between 3 and 15 times. The volume the combustion chamber 10 is sufficient to substantially complete the combustion reaction, typically about a second.

As described in FIG. 2, the atomization-heating nozzle 2 preferably has a conical convergent-divergent nozzle 5 with the residual oil feed 1 entering along the axis of the conical convergent-divergent nozzle 5 as shown in the figure. The convergent-divergent nozzle may also be a conventional axisymmetric convergent-divergent nozzle with the residual oil feed 1 entering through its periphery. An array of axisymmetric convergent-divergent nozzles may be used. The preferred conical convergent-divergent nozzle 5 can be conveniently formed between an inner atomization-heating nozzle body 12 and an outer atomization-heating nozzle body 13. The temperature of both the inner atomization-heating nozzle body 12 and the outer atomization-heating nozzle body 13 is controlled with a nozzle coolant stream 15. The nozzle coolant stream 15 may be a liquid phase coolant, water or Dowtherm A for example. The nozzle coolant stream 15 may advantageously contain both liquid and vapor phases to improve temperature control, increase the dimensional stability of the inner and outer inner atomization-heating nozzle bodies, and, in turn, the dimensional stability of the conical convergent-divergent nozzle 5. The two phase coolant pressure is advantageously set to maintain a coolant temperature between 100 and 350° C., more preferably between 200 and 300° C. The nozzle coolant 15 is intimately contacted with the inner 12 and outer body 13 of the atomization-heating nozzle 2 in annular coolant feed 16 and return 17 channels. The feed 16 and return 17 channels can be advantageously fitted with helical baffles to increase the velocity of the coolant and to ensure that the coolant flow is uniformly distributed.

The combustion gas 4 enters the atomization-heating nozzle 2 through gas feed port 19 and is distributed by an annular header 20 to the inlet of the annular convergent-divergent nozzle 5. The combustion gas 4 is accelerated by the conical convergent-divergent nozzle 5 to produce the combustion gas jet 3. The angle 21 between the combustion gas jet 3 and the residual oil feed 1 is preferably between 30 and 75 degrees, more preferably between 45 and 60 degrees. The ratio of the combustion gas jet 3 to resid feed 1 flow rate is adjusted to obtain an adiabatic temperature for the reactant stream 6 between 500 and 850° C. A conduit 22 connects the atomization-heating nozzle 2 with an inertial vapor-liquid separation device 23. The preferred inertial separation device 23 is a cyclone with several additional features. Methods are readily available (Green-1999) to determine cyclone design parameters and estimate cyclone separation performance and residence time as a function of the cyclone operating conditions and design. Readily available thermal cracking kinetic models (Gray-2004) can be used to estimate the extent of thermal degradation as of function of operating conditions.

As further described in FIGS. 3, the inertial device feed 24 tangentially enters the cylindrical section 25 of the inertial device 23 through a rectangular conduit 26. The vapor components 27 follow a helical flow pattern to the vapor discharge cylinder 28. The vapor components 27 are rapidly cooled to a temperature less than 400° C., more preferably less than 300° C., by intimate contact with a liquid quench stream 29 to form the vapor product stream 30. A conventional spray nozzle 30 is used to increase the liquid quench interfacial area and distribute the small quench liquid droplets across the cross section of the cylinder vapor discharge conduit 28. The atomizer may use the liquid quench stream 29 as the motive fluid or a vapor motive fluid. Steam or vaporized quench fluid 29 are convenient vapor motive fluids. The normal boiling point of the liquid distillate quench stream 29 is preferably between than 25 and 150° C. less than the light oil 31 operating temperature, more preferably between 50 and 100° C. less than the light 31 operating temperature.

The liquid components in stream 24 follow a helical flow pattern downward through the cylindrical 25 and tapered 32 sections to the heavy oil discharge 33. The tapered 32 section uses inertial forces to accelerate the heavy oil to the heavy oil discharge 33 conduit and increase the heavy oil film thickness at the heavy oil discharge 33 conduit inlet. The tapered section 32 typically uses a straight line taper with the difference between the radius at the bottom and the radius at the top of the tapered section 32 is roughly equal to the heavy oil discharge 33 conduit width 40.

The cylindrical 25 and tapered 32 section walls are advantageously cooled to a temperature between 400 and 200° C., more preferably between 350 and 250° C. using a cooling jacket 35. A process derived hydrocarbon stream or heat transfer fluid, e.g. Dowtherm A, may be conveniently used as cyclone jacket coolant 34. The jacket coolant is preferably a narrow boiling range hydrocarbon mixture with its bubble point essentially equal to the desired wall temperature at a convenient jacket operating pressure to improve the temperature control of the cylindrical 25 and conical 32 section walls. The hydrocarbon vapor in the coolant return stream 36 may be condensed to produce steam and regenerate the jacket coolant feed stream 34. A liquid distillate quench 39 may be advantageously added to cylindrical 25 or tapered 32 sections or heavy oil discharge conduit 33 (as shown on FIG. 3) of the cyclone 23 to accelerate cooling of the liquid portion of the cyclone feed 24. A purge gas 37 may be advantageously fed to the cyclone tapered section 32 (as shown on FIG. 3) to purge the vapor components of the cyclone feed 24 and minimize condensation on the cooler cylindrical 25 and tapered 32 section walls of the cyclone. In addition, the thickness of the heavy oil film on the tapered 32 section walls can be control by controlling the rate of vapor exiting via the heavy oil discharge conduit 33. The preferred purge gas is steam,

The equipment in FIGS. 1, 2, and 3 may be operated in three modes: flash deasphalting, flash pyrolysis, flash H₂ donor cracking. The process objective for a flash deasphalting unit operation is to produce the desired quantity of asphaltic heavy oil from the carbonaceous residual oil feed by primarily flash vaporization with minimum thermal degradation of both the light and heavy oil products. This objective is achieved by operating at the minimum residence time and temperature required to achieve the desired separation. Since this process has a very low yield of low boiling point hydrocarbons, air is the preferred oxidant.

The process objective of the flash pyrolysis unit operation is to increase the operating temperature to achieve higher light oil yields by a combination of vaporization and thermal cracking of the heavy oil. Since this process produces a significant yield of low boiling point hydrocarbons, a substantially pure oxygen oxidant with steam dilution is the preferred oxidant.

The light oil yield can be further increased by hydrotreating the residual oil feed with a homogeneous catalyst to more effectively hydrogenate the asphaltene species to produce a hydrogen donor solvent. The entrained homogenous hydrogenation catalyst from the hydrogenation step also catalyzes the hydrogen transfer reactions in the flash process to further decrease the rate of coke precursor formation. The light oil yield can be further increased by using a residual oil pretreatment method comprising adding colloidal molybdenum sulfide catalyst to a carbonaceous residual oil, recycling a portion of the heavy oil product, combining the residual oil feed and recycle carbonaceous heavy oil, feeding this mixture to a hydrotreater to add between 100 and 1500 standard cubic foot (SCF) H₂/bbl, more preferably between 150 and 1000 SCF H₂/bbl, most preferably between 200 and 700 SCF H₂/bbl. The reactant residence time above 425° C. is preferably less than 50 milliseconds and the reactant temperature is set to provide a conversion of asphaltic species to heptane soluble species of less than 85%, more preferably less than 75%.

Example 1

This example illustrates the use of the flash processing method to produce a deasphalted oil product from a typical oil sand bitumen and with sufficient by-product steam for bitumen production using the SAGD process. In this example, the oil sand bitumen feed rate is 33.5 metric tons per hour. The combustion chamber fuel requirement is 1.7 metric tons per hour of the light oil product. The combustion chamber was operated at a pressure of 5.3 bar. The combustion gas has a temperature of 2000° C. with a 10% oxygen deficiency. The conical convergent divergent nozzle has a residual oil feed—gas jet angle of 60 degrees and a throat area of 130 cm². The reactants have an average temperature of about 540° C. and a pressure of 2.2 bar. The cyclone separator has a 14×29 cm feed conduit, cylindrical section diameter of 58 cm and height of 36 cm, gas exit diameter of 29 cm and height of 36 cm, a pressure atomizer with quench liquid flow rate of 24.4 metric tons per hour, light oil temperature of 400° C., a tapered section minimum diameter of 58 cm, maximum diameter of 65 cm and height of 80 cm, a heavy oil discharge conduit width of 3.5 cm and height of 7 cm, a stripping steam flow rate of 9.4 metric tons per hour, a heavy oil product rate of 5.7 metric tons per hour (roughly equivalent to the asphaltene content of the oil sand bitumen feed). Thermal cracking reactions resulted in about 1% of the asphaltenes converted to lighter products. Combustion of the heavy oil provides a net steam production of about 2.65 kilograms per kilogram oil sand bitumen feed, which is a typical steam requirement for oil sand production by the SAGD process. The deasphalted oil product flow rate is about 26.2 metric tons per hour.

Example 2

Example 2 uses the equipment described in the first Example 1 to convert asphaltenes species to deasphalted oil and heavy distillates with minimum production of gaseous and coke precursor species. 23 kilograms per hour of a 535° C.+oil sand vacuum resid, with an asphaltene convent of about 33 wt %, is dosed with 1000 ppm of a molybdenum sulfide colloidal catalyst and blended with 10.4 kilograms/hour of heavy oil recycle, hydrotreated in an ebullated hydrotreater at 100 bar for sufficient liquid residence time to add roughly 640 SCF/bbl of hydrogen to the resid. 3 kilograms/hour of a natural gas fuel are oxidized at 5.3 bar pressure with an oxidant comprising 11.3 kilograms/hour of substantially pure oxygen with a steam diluent to oxygen weight ratio of 1.2:1. The resulting combustion gas is expanded from 5.3 bar to 2.2 bar through a conical convergent-divergent nozzle and contacted with the partially hydrogenated vacuum resid and recycle heavy oil. This procedure achieves a maximum thermal cracking temperature of roughly 700° C. and results in the conversion of about 93% of the asphaltenes in the vacuum resid feed. The unconverted asphaltic heavy oil by-product contains substantially all the metals in the vacuum resid and catalyst feeds in addition to the coke precursors and any coke produced by the process. This material is preferably fed to a gasification unit to produce hydrogen and recover the metal values.

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 upgrading asphaltic residual oil comprising:

a) producing a combustion gas jet by expansion of a combustion gas through a convergent-divergent nozzle

b) contacting said asphaltic residual oil with said combustion gas jet;

c) separating a vapor phase and a liquid phase formed by said contacting; and

d) cooling said vapor phase and said liquid phase products to form light and heavy oil products.

2. The method as claimed in claim 1 wherein said cooling is to a temperature less than 400° C.

3. The method as claimed in claim 1 wherein said asphaltic residual oil feed has an apparent viscosity less than 1000 cP.

4. The method as claimed in claim 1 wherein said asphaltic residual oil is at a temperature of about 100 to about 425° C. prior to contacting said combustion gas jet.

5. The method as claimed in claim 1 further comprising hydrotreating said asphaltic residual oil by contacting with a homogeneous catalyst prior to contact with said combustion gas.

6. The method as claimed in claim 1 wherein said homogeneous catalyst is a molybdenum sulfide catalyst.

7. The method as claimed in claim 1 wherein said hydrotreating is performed in an ebullated or fixed bed hydrotreating reactor.

8. The method as claimed in claim 1 wherein hydrogen is added to said asphaltic residual oil.

9. The method as claimed in claim 1 wherein said hydrogen is added to said asphaltic residual oil at about 100 to about 1500 standard cubic feed per barrel of asphaltic residual oil.

10. The method as claimed in claim 1 wherein said hydrogen is added to said asphaltic residual oil at about 150 to about 1000 standard cubic feed per barrel of asphaltic residual oil.

11. The method as claimed in claim 1 wherein said combustion gas jet is formed by reacting an oxidant with a fuel.

12. The method as claimed in claim 1 wherein said oxidant is selected from the group consisting of air, oxygen-enriched air and substantially pure oxygen.

13. The method as claimed in claim 1 wherein said fuel is selected from the group consisting of carbon monoxide, hydrogen, gaseous hydrocarbons and mixtures thereof.

14. The method as claimed in claim 1 further comprising adding steam to said fuel.

15. The method as claimed in claim 1 wherein the pressure ratio of said combustion gas to said reactant gas stream is between 2 and 20.

16. The method as claimed in claim 1 wherein the temperature of said combustion gas jet is a bout 1250 to about 2000° C.

17. The method as claimed in claim 1 wherein said convergent-divergent nozzle is axisymmetric.

18. The method as claimed in claim 1 further comprising an array of convergent-divergent nozzles.

19. The method as claimed in claim 1 wherein said asphaltic residual oil contacts said combustion gas jet from the periphery of said combustion gas jet.

20. The method as claimed in claim 1 wherein said convergent-divergent nozzle is conical.

21. The method as claimed in claim 1 wherein said asphaltic residual oil contacts said combustion gas jet along the axis of said conical convergent-divergent nozzle.

22. The method as claimed in claim 1 wherein said separation is performed in a cyclone.

23. The method as claimed in claim 1 wherein said reactant residence time above 425° C. is less than 400 milliseconds.

24. The method as claimed in claim 1 wherein a liquid distillate quench is used to rapidly cool said vapor phase and said gas phase.

25. The method as claimed in claim 1 wherein said quench is atomized.

26. The method as claimed in claim 1 wherein a purge gas is used to minimize loss of said liquid distillate in said heavy oil product.

27. The method as claimed in claim 1 wherein ratio of said combustion gas jet and said asphaltic residual feed flow rate is adjusted to obtain an adiabatic temperature for said reactant stream between 500 and 850° C.