CA1298803C - Viscosity reduction by direct oxidative heating - Google Patents

Abstract

ABSTRACT
A method is disclosed for reducing the viscosity of a hydrocarbon feed. The feed is heated from an ini-tial temperature to a second temperature and an oxidiz-ing agent is introduced to oxidize components in the feed and provide heat to increase the temperature of the feed to a reaction temperature. The reaction tem-perature is maintained to produce a reaction product having a lower viscosity than the feed.

Description

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IMPROVED VISCOSITY REDUCTION
BY DIRECT OXIDATIVE HEATING

Field of the Invention This invention relates to a method for improving 05 the transportability of heavy oils and other hydrocar-bons by thermal viscosity reduction with reduced coke formation on reactor walls wherein an incremental por-tion of the heat is provided by direct oxidative heat-ing of the hydrocarbon material.

Background of the Invention Vertical tube reactors which ordinarily involve the use of a subterranean U-tube configuration for providing a hydrostatic column of fluid sufficient to provide a sPlected pressure are well known. This type of reactor has been primarily used for the direct wet oxidation of materials in a waste stream and par-ticularly for the direct wet oxidation of sewage sludge. Bower in U.S. Patent No. 3,449,247 discloses a process in which combustible materials are disposed of by wet oxidation. A mixture of air, water and combus-tible material is directed into a shaft and air is in-jected into the mixture at the bottom of the hydros-tatic column.

2~ Lawless in U.S. Patent No. 3,606,999 discloses a similar process in which a wate~ solution or suspension of combustible solids is contacted with an oxygen-containing gas. Excess heat is removed from the ap-paratus by either diluting the feed with the product 3~ stream or withdrawing vapor, such as steam, from the system.
Land, et al. in U.S. Patent No. 3,464,885 (issued September 2, 1969) is directed to the use of a subter-ranean reactor for the digestion of wood chips. The method involves flowing the material through counter-current coaxial flow paths within a well bore while ' .. . . . . . . ...

3 `~-flowing heated fluid coaxially of the material to be reacted. The reactants, such as sodium hydroxide and sodium sulfate, are combined with the wood chip stream prior to entry into the U-tube which is disposed within a well bore.
05 Titmas in U.S. Patent No. 3,853,759 (issued Decem-ber 1~, 1974) discloses a process in which sewage is thermally treated by limiting combustion of the material by restricting the process to the oxygen which is present in the sewage, i.e. no additional oxygen is added. Therefore, it is necessary to provide a con-tinuous supply of heat energy to affect the thermal reactions.
McGrew in U.S. Patent No. 4,272,383 (issued June 9, 1981) discloses the use of a vertical tube reactor to contact two reactants in a reaction zone. The method is primarily directed to the wet oxidation of sewage sludge in which substantially all of the organic material is oxidized. Heat exchange between the in-flowing and product streams is contemplated. The tem-perature in the reaction zone is controlled by addingheat or cooling as necessary to maintain the selected temperature. It is disclosed that when gas is used in the reaction, it is preferred to use a series of en-larged bubbles known as "Taylor bubbles". These bubbles are formed in the influent stream and passed downward into the reaction zone. It is disclosed that preferably air is introduced into the influent stream at different points with the amount of air equaling one volume of air per volume of liguid at each injection point. While such a large amount of oxygen can be needed to oxidize minor organic components dissolved o_ suspended in a primarily aqueous liguid, this process is not feasible when the li~uid stream is primarily a mixture of hydrocarbons. The presence of such large -~ ~Z~ 3 -~

volumes of oxygen could result in an uncontrollable ex-othermic reaction.
T~e above-cited patents w~ich disclose vertical tube reactor systems describe the use of such systems with primarily aqueous streams. None of these patents 05 describe treatmen, of a primarily ~ydrocarbon stream.
Specifically, there is no suggestion of the t~ermal treatment ~f a ~ydrocarbon stream in a vertical tube reactor system.
T~e reduction in viscosity of heavy hydrocarbon material by t~ermal treatment are well known. The thermal cracking known as rvisbreaking" involves the treatment Df hydrocarbon materials at elevated tempera-tures and pressures. Such processes are exemplified by Biceroglu, et al. in U.S. Patent No. 4,462,~9~ (1984), 1~ Beuther, et al. in U.S. Patent No. 3,132,D88 (1964), ~aff, et al. in U.S. Patent No. 2,695,264 (19~4), and Shu, et al. in U.S. Patent No. 4,504,377 (198~). Such processes are commonly used in refineries where there are the necessary distillation units to provide selec-2~ tive fractions to the visbreaking u~it and the neces-sary product treatment facilities to handle the gaseous ~nd low boiling products from the visbreaking unit.
Such capital intensive processes do not readily lend themselves to t~e treatment of heavy ~ils at the production site to improve their transportability.
Applicant~ 6 own United State6 Patent No. 4,778,586, issued October 18, 1988 di6close6 a method for viscosity reduction of a hydrocarbon feed in the field. In this proce6s a vertical tube reactor is used to create a hydro6tatic pressure on the crude oil feed and the feed is heated by an external heat source to provide the viscosity reduction ~ece6sary to improve transportability of the feed from the production area. The temperature differential between the heat 8 ource 3~

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lZ9~3 and the feed is maintained small to minimize the forma-tion of coke.
Commonly assigned U.S. Patent No. 4,648,964 of Leto et al. (1987) discloses the use of a vertical tube reactor to separate hydrocarbons from tar sands froth.
~ The formation of coke deposits on the walls of the reaction vessels or heating surfaces has been a con-tinuing problem. It has been disclosed that at higher severities there is an increased tendency to form coke deposits in the heating zone or furnace. Black in U.S.
Patent No. 1,720,070 teaches that operating at lower temperatures for increased lengths of time provides `'a much smaller amount of carbon is deposited than is deposited at higher temperatures." Akbar et al., "Visbreaking Uses Soaker Drum", Hydrocarbon Processing, 1~ May 1981, p. 81 discloses that, when there is a high temperature differential between the tube wall in a furnace cracker and the bulk temperature of the oil, the material in the boundary layer adjacent to the tube wall gets overcracked and excessive coke formation oc-curs. In furnace cracking this boundary layer is com-monly about 30C to 40C higher than the bulk tempera-ture.
The problem associated with excessive coke forma-tion in the boundary layer stems from the fact that the coke adheres to vessel walls. This coating of material acts to insulate the reaction vessel which necessitates additional heating for sufficient viscosity reduction.
The added heat compounds the problem by further in-creasing coke formation.
In refinery operations, coke formation in vis-cosity reduction processes can be tolerated because frequent shutdowns of the process for coke removal are possible since storage space for the feedstock is usually available. However, this limitation is unac-3~ ceptable in a field operation where crude is con-. .. . . . .

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tinually produced and must be rapidly transported.
Such periodic shutdowns are als~ unacceptable with a vertical tube reactor system. I~ applicant's own u.S~
U.S. patent No 4,778,586, the temperature dif-ference between the heat source and the feed is kept Q~ small to minimize formation of coke. However, this process still has the limitation that the temperature of the wall of the reaction vessel is necessarily higher than the temperature of the bulk of the hydrocarbon stream. Consequently, over a period of time coke formation can occur which re~uires either a decoking operation or shutdown of the unit.
Accordingly, there is a need for an improved method for reducing the viscosity of recovered heavy hydrocarbon material in which coking of reactor vessels can be substantially reduced.
The ~resent invention provides-a method for reduc- ~ -ing the viscosity of a hydrocarbon feed in which a final incremental amount of heat necessary for in-creased thermal degradation of heavy comp~nents is provided by the exothermic oxidation of components in the feed. This process avoids undesirable coking in the reactor vessel by maintaining the temperature in the boundary layer of the stream near the vessel walls below coking temperatures.
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Summarv of the Invention The present invention comprises a process for reducing the viscosity of a hydrocarbon composition in which a feed stream of the composition having a core 3D portion and a boundary layer is introduced into a ves-sel. The bulk temperature of the stream is increased from a first bulk temperature to a second ~ulk tempera-ture. An oxidizing agent is introduced into the core portion of the stream to oxidize components in the stream and provide heat to the core portion of the _~_ . . . .

t~3 stream to provide a bulk reaction temperature greater than the second temperature. The amount of the oxidiz-ing agent is controlled to maintain the reaction tem-perature below the coking temperature of the feed. The reaction bulk temperature is maintained to produce a 05 reaction product having a lower viscosity than the feed~
In another embodiment, the instant invention com-prises a method for reducing viscosity of a hydrocarbon composition using a vertical tube reactor. An influent stream of the hydrocarbon feed is increased from a first temperature to a second temperature by heat ex-change between the influent stream and effluent product stream. At least one of the streams is in turbulent flow during the heat exchange. The pressure on the hydrocarbon feed is increased from a first pressure to a second pressure by a hydrostatic head. An incremen-tal amount of heat necessary to increase the bulk tem-perature of the feed from the second temperature to a reaction temperature is provided by introducing an oxidizing agent into the core portion of the feed stream to o~idize components in the feed.
In another embodiment, the instant invention com-prises a method for reducing the viscosity of a hydrocarbon feed by thermal degradation of heavy molecular weight components of the feed at a reaction temperature. The feed is heated with a heat source to below a reaction temperature. The incremental amount of heat necessa~y to heat the feed to the reaction tem-perature is provided by internal combustion of a por-tion of the feed.

-~-~ Brief Descriptlon of the Drawin~
FIG. 1 is a schematic representation of apparatus useful in the practice of the present process; and FIG. 2 is a representation of a preferred method of operation of the instant process.

Detailed Description of the Invention As used herein, the term "boundary layer" is 05 defined as the thin layer of the hydrocarbon stream im-mediately adjacent to reactor walls or other stationary surfaces in the reactor vessel, this layer being characterized by very low fluid velocities.
As used heI~in, the term "core portion" is defined as the portion of the hydrocarbon stream other than the boundary layer which is characterized by flow velocities which are higher than boundary layer flow velocities. The core portion can be in laminar or tur-bulent flow.
As used herein, the term "bulk temperature" is defined as the average temperature in a cross-sectional segment of the core portion in the hydrocarbon stream in which there is sufficient mixing of the stream to achieve a substantially uniform temperature throughout the segment.
As used herein, the term "coking temperature" is defined as a bulk temperature at which there is at least about 0.5 weight percent solid coke formation in a 24 hour period tbased on the hydrocarbon stream).
The present invention involves providing an in-cremental amount of heat to a hydrocarbon stream by in-troducing an oxidizing agent into the core portion of the stream. The oxidizing agent rapidly oxidizes com-ponents in the stream in an exothermic oxidation reac-tion. By distributing this heat in the moving stream, an increase in the bulk temperature of the stream is provided. This reaction temperature is the temperature at which the rate of viscosity reduction is substan-tially increased. The oxidation reaction is controlled so that the increased ~ulk temperature (reaction 129~ 3 temperature) is below the coking temperature. As dis-cussed above, maintaining the bulk temperature below the coking temperature limits the temperature of the boundary layer in the reactor vessel which prevents ex-cessive formation of coke on the walls of the reactor 05 vessel.
It has been found that by practice of the present invention, the viscosity of a hydrocarbon feed can be significantly reduced without the formation of sub-stantial coke deposits on the walls of the reactor ves-sel. While the process of coking is not fully under-stood, it has been reported that increased severity of conditions increases coke formation. It is known that materials such as asphaltenes are more likely to form coke. Once these materials precipitate and solidify on surfaces, it is difficult to dissolve them before coke deposits are formed. Coke tends to build on the reac-tor wall or other heating surface because in most sys-tems these surfaces must be heated significantly above the desired reaction temperature to attain bulk tem-peratures sufficient to effect acceptable rates of ~is-cosity reduction. Such "external heating" promotes coke formation on reactor walls.
Practice of the present invention avoids these problems associated with external heating. The incre-ment of heat necessary to increase the bulk temperatureof the stream to effect substantially increased rates of viscosity reduction is provided by internal heating through direct oxidation of components in the core por-tion of the stream. Consequently, coke formation on reactor walls or other-surfaces in the reactor vessel is substantially reduced since these surfaces and the boundary layer of feed adjacent to the surfaces are not heated above the coking temperature.
While practice of the present invention substan-tially reduces formation of coke on reactor vessel ~ ~ 2~ 03 walls, some coke formation can occur over time. Theamount of coke build-up is affected by the type of feed, the quantity of feed which is processed as well as process conditions. While some coke build-up can be tolerated in most viscosity reduction processes, the 0~ present invention is less sensitive to coke formation than systems which rely entirely on external heating.
Coke fo~mation on reactor walls insulates the reactor and decreases the amount of heat added to the stream by an external heat source. To maintain required tempera-tures for viscosity reduction, external heat must beincreased which causes additional coke formation.
However there is a significant advantage in the present process since coke formation in the reactor does not require additional external heating because the final increment of heat is provided internally. The amount of coke formation in the present process which would necessitate a decoking procedure depends on the par-ticular reaction vessel in use and the point at which the operation becomes impaired by coke buildup.
Internal heating is achieved by oxidizing a part of the core portion of the hydrocarbon stream. This exothermic reaction is controlled so that the bulk tem-perature remains below the coking temperature. It should be appreciated that between the region in the reactor vessel where the oxidation reaction occurs and where mixing of the stream has achieved a substantially uniform temperature throughout a cross-sectional seg-ment of the stream, localized temperatures above the coking temperature can be expected to occur. Such tem-peratures can cause some coke formation in the stream.These coke particles, however, can be substantially prevented from adhering to any surfaces by the physical action of the flow of the stream.
It was anticipated that direct oxidation of the hydrocarbon stream would cause formation of oxygenated . .

12~3~803 ~

by-products, such as aldehydes, ketones or carboxylic acids. Surprisingly, it has been found that production of these and similar components by the present process is unexpectedly low. This result is beneficial because the presence of CUch compounds lowers the value of the 05 hydrocarbon product and can result in decreased storage stability of the product. It has been unexpectedly found that the primary products of the oxidation reac-tion are carbon dioxide, carbon monoxide and water.
The process of the present invention is broadly applicable to reducing the viscosity of hydrocarbon feeds. The terms "hydrocarbon stream" and "hydrocarbon feed" are used interchangeably herein to mean a liquid stream which contains primarily hydrocarbonaceous com-ponents but can also contain smaller amounts of other components, for example, water. The present invention is especially useful for treating heavy oil crudes of a nature and viscosity which renders them unsuitable for direct pipeline transport. This includes feeds having a viscosity above about 1000 centipoise (cp) at 25C
(unless otherwise indicated, viscosity referred to herein is at 25C), a pour point above about 15C or an API gravity at 25C of about 15 and below. The advan-tages of reduced viscosity, increased API- gravity and/or reduced pour point can be achieved by practice 2~ of the present invention without regard to the initial viscosity, API gravity or pour point of the feed. Ad-ditionally, if desired, a diluent can be added to the feed stream or to the reaction product from the instant process in order to further reduce the viscosity.
Heating of the product in order to reduce the viscosity or maintain an acceptable viscosity for a particular pipeline or transportation medium is also possible.
Hydrocarbon feeds which can be used in the instant process include, but are not limited to, heavy whole crude oil, tarsands, bitumen, kerogen, and shale oils.

~xamples of heavy crude oil are Venezuelan Boscan crude oil, Canadian Cold Lake crude oil, Venezuelan Cerro Negro crude oil and California Huntington Beach crude oil. In practice, the most significant reductions in viscosity are achieved when the starting feed is more 05 viscous.
The vertical tube reactor system useful in the in-stant invention has a heat exchange section, a combus-tion zone, and a reaction zone. The heat exchange sec-tion is adapted to provide for heat exchange between the influent hydrocarbon feed stream and the effluent product stream. The combustion zone is the region in which oxidizing agent is introduced into the core por-tion of the hydrocarbon stream. The reaction zone is the region in which the bulk temperature of the hydrocarbon stream is greater than the maximum tempera-ture achieved by heat exchange. There can be substan-tial overlap between the combustion zone and the reac-tion zone.
In the instant process, the hydrocarbon feed stream comprising a core portion and a boundary layer is introduced into the inlet of the vertical tube reac-tor. The influent hydrocarbon stream is at a first temperature (Tl) and an initial pressure (P1). As the influent hydrocarbon stream travels down the vertical 2~ tube reactor, the pressure increases due to the hydros-tatic column of fluid. Additionally, the temperature of the influent stream increases to a second tempera-ture (T2) due to heat exchange with the effluent product stream. An oxidizing agent is introduced into the core portion of the hydrocarbon stream to increase the bulk temperature of the hydrocarbon stream to a pre-selected reaction temperature (TrX).
It is important that the temperature increment be-tween the second temperature and the reaction temper~-ture is small because less feed must be consumed in the . .

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oxidation reaction to provide the necessary heat andfewer oxidation products are formed. Additionally, the greater the temperature increment, the larger the com-bustion zone needed to provide the necessary heat to increase the bulk temperature of the stream from the 05 second temperature to the reaction temperature. It is preferred that the temperature increment between the reaction temperature and the second temperature of the hydrocarbon stream is less than about 35C and more preferably less than about 25C.
In order to achieve the second temperature neces-sary for the instant process to operate efficiently, it is necessary for the heat exchange between the influent hydrocarbon stream and the effluent product stream to be more efficient than those disclosed in the known patents relating to vertical tube reactors. The tem-perature of the influent stream achievable by heat ex-change with the reaction product is limited by a number of factors including the temperature of the reaction product, the heat exchange surface area, and the velocities of the hydrocarbon streams. In order to achieve the necessary heat exchange efficiencies, it has been found that at least one of and preferably both the influent feed stream and the product stream are in substantially vertical multiphase flow. It has been found that when both streams are in substantially ver-tical multiphase flow an increase in heat exchange ef-ficiency of at least about 100~ can be achieved com-pared to heat exchange when neither stream is in multi-phase flow. This allows a second temperature to be at-tained which is sufficiently close to the necessaryreaction temperature to allow direct oxidative heating by introducing an oxidizing ayent.
The oxidizing agent of the present invention is a material which rapidly exothermically oxidizes the hydrocarbon feed under chosen reaction conditions. The Q;~ ~

agent is selected so that essentially all of the agent reacts with the feed. Various oxidizing agents are suitable for use in the present invention. Such agents include, but are not limited to oxygen and hydrogen peroxide. The oxidizing agent can be optionally mixed 05 with a nonreactive gas, such as nitrogen, and air or enriched air can be used in the present process.
Preferably enriched air is used.
The amount of the oxidizing agent injected into the hydrocarbon stream affects the amount o~ heat gen-erated by the oxidation reaction and is the primaryfactor for controlling the temperature increase in the stream from the oxidation reaction. The amount of oxidizing agent required for a particular volume of hydrocarbon feed in operation of the invention can be substantially defined with four variables: (l) the heat required to raise the temperature of that volume of the feed from the second temperature to a reaction tempera-ture, (2) the heat of cracking of that volume of the feed, (3) the heat loss from that volume of the feed to the environment in the reaction zone, and (4) the heat of combustion of the particular feed. The sum of the first three of these guantities equal the amount of heat that must be generated from the oxidation of some portion of the feed. The amount of feed which must be oxidized depends on the heat of combustion of the par-ticular feed.
With regard to the variables discussed above, it is apparent that as the difference between the second temperature and the reaction temperature increases an increased flow rate of oxidizing agent is necessary to generate additional heat by the oxidation of a larger amount of the feed. As stated above, the amount of oxidizing agent required in the process is also depen-dent on the heat of cracking of the feed. This charac-teristic is variable between feeds. The oxidizing .. . ..

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agent ~low rate is also affected by heat loss from thehydrocarbon stream to the environment. A greater heat loss requires more heat generation initially and, therefore, the use of more oxidizing agent.
In operation of the invention, the amount of 05 oxidizing agent introduced to the reactor vessel is used to control the oxidation reaction. The desired flow rate for a given concentration can be estimated by calculation using the variables discussed above. If the exact values for each variable is known, the amount of oxidizing agent required (assuming the heat of oxidation is known) can be determined. In practice, these values must ordinarily be estimated. Such an es-timate can be used to determine an initial flow rate of oxidizing agent to use; however, actual control is based on a measured variable such as the bulk tempera-ture of the hydrocarbon stream. The bulk temperature downstream from the oxidation reaction is ordinarily monitored. The bulk temperature should remain below the coking temperature so that the reactor walls and boundary layer are not heated to a temperature where excessive coke formation occurs. If the bulk tempera-ture becomes too high, the flow of oxidizing agent is reduced until the preselected bulk temperature is at-tained. If the bulk temperature is too low to achieve 2~ acceptable viscosity reduction, the amount of oxidizing agent introduced into the stream is increased until the appropriate reaction temperature is attained. Monitor-ing the pressure in the reaction zone can also be used to control the amount of-oxidizing agent introduced into the hydrocarbon stream. The detection of pressure surges or fluctuations indicates that the amount of oxidizing agent beiny introduced into the hydrocarbon stream should be decreased.
As used herein, the term "reaction temperature"
refers to the maximum bulk temperature of the hydrocar-^~ 12~ 3 -~

bon stream raachad in the process. It is understood that some thermal cracking can occur at lower tempera-tures. The term "reaction zone r refers to the region in the process which begins at the point the oxidizing ayent is introduced and ends where heat exchanye be-05 tween the reaction product effluent stream and the in-fluent hydrocarbon stream beyins. The maximum useful bulk temperature in the instant process is the coking temperature of the particular feedstock. In ordinary operation, the bulk temperature of the hydrocarbon stream is maintained below the coking temperature. At a minimum, the reaction temperature used for practice of the instant process is high enough to initiate some thermal cracking reaction. For most feeds, the reac-tion temperature is above about 300C and less than about 475C, more typically in the range of about 350C
to about 4~0C, and more often in the range of about 3~C to about 435C.
The hydrocarbon stream and reaction zone is preferably maintained under a superatmospheric pressure typically above about 1,0~0 pounds per square inch ab-solute (psi). The high pressure serves to maintain volatile components in the hydrocarbon stream in liquid phase. The pressure also maintains products and by-products from the oxidation reaction and thermal crack-iny reaction in solution in the hydrocarbon stream. Itis important to maximize the li~uid phase in the reac-tion zone to minimize the concentration of asphaltenes and other coke precursors to avoid their precipitation from the hydrocarbon phase and possible deposition on internal reactor surfaces with subsequent coke forma-tion. A small volume fraction of the stream can be in vapor phase and, in fact, a small volume of vapor phase can be beneficial in promoting mixing of the stream for rapid distribution of heat from the oxidation reaction throughout the core portion of the stream. Preferably --1~--$~3 the vapor phase should amount to no more than about lO
volume percent of the hydrocarbon stream. If the vapor phase comprises a substantial percent of the stream volu~e, it can become difficult to maintain a pressure balance in th* reactor vessel.
05 AS discussed hereinabove, at least a portion of the pressure on the hydrocarbon stream is achieved by a hydrostatic column of fluid. If it is desired that the reaction pressure be greater than that generated by the hydrostatic head, the initial pressure of the hydrocar-bon feed stream can be increased by, for example, centrifugal pumps, to provide the desired total reac-tion pressure.
Vpon introduction of the oxidizing agent into the hydrocarbon stream, oxidation of components of the stream occurs upon contact with the oxidizing agent.
In a localized area immediately downstream from intro-- duction of the agent, the temperature of the stream can be substantially higher than the reaction temperature because- the---oxidation-reaction~occurs essentially upon contact of the agent with hydrocarbon materials and is substantially complete before the heat generated by the reaction is dissipated in the stream. The use of oxygen as the oxidizing agent results in essentially a flame front in the hydrocarbon stream. It is desirable to very quickly distribute the heat from the oxidation reaction throughout the core portion to produce a sub-stantially uniform temperature in the core portion, i.e. essentially a uniform bulk temperature. Mixing of the core portion ordinarily occurs essentially im-mediately as a result of turbulent flow of thehydrocarbon stream within the reaction vessel. If the flow velocity of the stream is low enough that the stream is in laminar flow, mixing can be induced with, for example, static mixers.

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The ra~e at which the oxidizing agen~ is intro-duce~ into the hydrocarbon stream can be conveniently expressed as an amount of oxidizing agent per unit volume of the hydrocarbon stream. The flow rate of the oxidizing agent is controlled so that the heat gen-05 erated by the oxidation reaction does not increase thebulk temperature of the hydrocarbon stream above the coking temperature. For example, in a typical opera-tion in which the hydrocarbon stream comprises whole crude oil and oxygen is the oxidizing agent, the flow rate of oxygen is preferably less than about 40 scf/bbl (standard cubic feet per barrel), more preferably less than about 30 scf/bbl and most preferably less than about 20 scf/bbl.
The primary gaseous product of the oxidation reac-tion has been found to be carbon dioxide, which corre-lates closely with introduction of oxygen to the reac-tor. Other gases are also produced as by-products of the present process, however, these appear to correlate with temperature fluctuations in the stream rather than the combustion reaction. The ma;or component of this gas make has been found to be methane with smaller amounts of ethane, propane, hydrogen, carbon monoxide, and hydrogen sulfide also being produced.
In operation of the present invention, it is im-portant-to maintain a positive pressure at the point of introduction of the agent into the stream. Otherwise, the hydrocarbon feed can flow into the oxidizing agent feedline possibly resulting in a violent oxidation reaction. Safe operation of the present process there-fore, requires that the oxidizing agent be at a pres-sure greater than the pressure of the feed at the point of injection. To maintain a positive oxidizing agent flow and prevent the danger of hydrocarbon backup-into the oxidizing agent addition line, a pressure drop .. .... .....

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across the in;ection nozzle of at least about 50 psi, and more preferably about 100 psi is preferred.
For safety reasons, it is also preferred to provide an emergency system in the event of a mechani-cal failure in the injection system. Such an emer~ency 05 system floods the injection line with a non-reactive gas, such as nitrogen, during an injection system failure to prevent hydrocarbon material from entering the injection line and producing an explosive reaction with the oxidizing agent.
The spatial placement of the oxidizing agent in-jection nozzle can significantly affect the temperature of regions of the boundary layer as well as the reactor vessel wall. If the nozzle is placed within the core portion of the hydrocarbon stream close to the boundary 1~ layer, the resulting oxidation reaction can heat the boundary layer and the reactor vessel and cause sub-stantial coke formation on the vessel. Likewise, if the injection nozzle is placed centrally within the core portion of the hydrocarbon stream but is directed toward a reactor wall or other surface, the resulting reaction can overheat the boundary layer and reactor vessel. Another danger associated with placement of the oxidizing agent in;ection nozzle is that if the nozzle is too near the reactor vessel wall or is pointed toward the reactor vessel wall, the oxidation reaction can degrade or melt the wall causing a system failure. In operation of the process, the oxidizing agent injection nozzie is located centrally in the core portion of the hydrocarbon stream and is directed on a line substantially parallel to the flow of the hydrocarbon stream. This placement of the nozzle acts to localize the oxidation reaction within the core por-tion of the hydrocarbon stream away from the boundary layer, thereby minimizing the temperature in the bound-33 ary layer.

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The injection nozzle should also be oriented rela-tive to the flow of the hydrocarbon stream so that heat generated by the oxidation reaction is carried away from the nozzle to prevent thermal deyradation of the nozzle itself. Injection of the oxidizing agent in the 05 same direction as the flow of the hydrocarbon stream, given a sufficient flow rate, successfully removes heat from the nozzle.
Heat loss to the outside environment from the central portion of the stream outward is anticipated as heat is generated internally by direct oxidative heat-ing. Some heat loss can occur even if the reactor ves-sel is insulated. Consequently, it may be necessary to use multiple sites for introduction of oxidizing agents to provide sufficient heat for viscosity reduction or to maintain a given temperature for a longer time than possible with a single injection site. In this embodi-ment, the injection sites are spaced so that as the bulk temperature of the stream falls below a tempera-ture at which acceptable viscosity reduction is occur-ring, the stream passes another injection site toprovide additional heat.
The instant invention can be more readily under-stood after a brief description of a typical applica-tion. As will be understood by those skilled in the art, other apparatus and configurations can be used in the practice of the present invention.
FIG. 1 depicts a subterranean vertical reactor 10 disposed in a well bore 12. The term "vertical`' is used herein to mean that the tubular reactor is dis-posed toward the earth's center. It is contemplatedthat the tubular reactor can be oriented seve-al de-grees from true vertical, i.e. normally within about 10 degrees. Duriny operation, flow of the hydrocarbon stream can be in either direction. As depicted, flow of the untreated hydrocarbon feed stream- is-throuyh ~29~3803 line 13 and into downcomer 14 to the reaction zone 16 and up the concentric riser 18. This arrangement provides for heat exchange between the outgoing product stream and the incoming feed stream. During start up, untreated hydrocarbon feed is introduced into the ver-05 tical tube reactor system through feed inlet 13, theflow rate being controlled by a valve 20. The hydrocarbon feed stream passes through downcomer 14 into reaction zone 16 and up through concentric riser 18 exiting through discharge line 22. During this operation unless external heat is provided to the hydrocarbon feed stream, the initial temperature Tl is equal to the final heat exchange temperature T2 and is also equal to the maximum temperature in the reaction zone TrX (assuming no heat loss to the environment).
1~ In order to achieve the necessary temperature T2 at which oxidant can advantageously be introduced, heat is provided to the hydrocarbon stream through external heating. This can be provided by an above ground heat-ing means 24.-~ The necessary heat-can also be provided by an external heating means 26 surrounding the reac-tion zone. Preferably, external heating means 26 is a jacket surrounding the reaction zone through which a heat exchange fluid is passed through inlet line 27 and outlet line 28. In another configuration not shown, the downcomer 14 can also be jacketed to allow external heating of the hydrocarbon stream at this location in addition to or instead of heating the reaction zone.
Alternatively, the external heating means 26 can be used in conjunction with the above ground heating means 24 to provide the hydrocarbon feed stream at the desired temperature ~2 As the hydrocarbon s~ream passes down through downcomer 14, pressure on any par-ticular volu~e segment increases due to the hydrostatic column of fluid above any particular point in the 3a stream. The temperature of the hydrocarbon stream is ....

dete.rmined by temperature monitors 29 which can be lo-cated in the hydrocarbon stream throughout the vertical tube reactor system. Pressure monitors 30 can also be located throughout the vertical tube reactor system to monitor any pressure increases or fluctuations in the 05 fluid stream.
Once the desired temperature T2 has been attained by external heating of the hydrocarbon stream, oxidant is introduced through line 32 to provide the incremen-tal heat necessary to ~ad the desired reaction tem-perature. As depicted, the oxidant enters thedownflowing hydrocarbon stream through one or more nozzles 34. Flow rate of the oxidant is controlled by valve 36 which in turn can be controlled directly or indirectly by output from selected temperature monitors 29 and/or pressure monitors 30. If needed, additional oxidant injection nozzles 3~ can be provided downstream from the initial nozzles 34. Nozzles 38 can be ac-tivated as needed to provide additional heat to the hydrocarbon-stream by activating valve 40. As dis-cussed hereinabove, for safet reasons it is importantto maintain a positive pressure in line 32 relative to the pressure of the hydrocarbon at the injection nozzle. This prevents hydrocarbon~~feed from flowing into the oxidizing agent feed line possibly resulting in a violent oxidation reaction. Therefore, the oxidizing agent should be at a pressure greater than the pressure of the feed at the point of injection, preferably a source of a non-reactive gas such as nitrogen. Nitrogen can be introduced into line 32 through line 42 with the flow being controlled by valve 44. Ordinarily, in operation line 32 is purged with nitrogen prior to introduction of oxidizing agent. For safety reasons, an emergency system is provided in which valve 44 is activated and non-reactive gas intro-3~

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duced into line 32 in the event oxidant flow is inter-rupted.
When the desired reaction temperature has been at-tained, heat from the external heat source can be ter-minated. As used herein, the term "external heat n does 0~ not apply to the heat provided to the influent stream by thermal communication with the effluent product stream.
The temperature of the effluent product stream may be somewhat lower than the reaction temperature when it initially comes in heat exchange contact with the in-fluent stream due to some heat loss to the environment.
The temperature of the effluent product stream is con-tinually decreased by thermal communication with the influent stream until a final temperature (Tf) is at-tained as the effluent exits the reactor system.
The effluent hydrocarbon stream passes upwardthrough riser 18 and out of heat exchange contact with influent hydrocarbon feed stream and out through line 22. The product can pass to a separation means 46 in which carbon dioxide and other gases are separated from liquid product and a more volatile fraction of the hydrocarbon stream can also be segregated. If desired, volatile components usually boiling below about 40C
can be recycled through line 48 into the influent hydrocarbon feed stream. This can be done to induce vertical multiphase flow in the influent stream to sub-stantially increase the efficiency of heat exchange be-tween the influent and effluent streams. Alterna-tively, during start up when external heat is being supplied to increase the temperature of the hydrocarbon stream, the complete stream can be recycled through line 48 in order to minimize the total volume of hydrocarbon which must be heated by exte~nal means. In an option (not shown), the product stream can be brought into thermal communication with the influent ~ ~g`~)3 -~

stream above ground to provide a higher initial tem-perature of the influent stream. Alternatively, the product stream can be cooled by mixing with unreacted hydrocarbon to improve transportability.
FIG. 2 depicts a preferred method of operation in 05 which the flow of influent feed is into the internal conduit 50 and up the external conduit 55. The initial nozzles 34 are located near the bottom of reaction zone 16. The nozzles are oriented to provide flow of oxidant essentially parallel to the flow of the feed stream. Additional nozzles 38 can be located down-stream from the initial nozzles. In operation, untreated feed passes down conduit 50 and product passes up through conduit 55. This method of operation has the advantage that vapor phase regions readily flow - 15 upward with the product stream. This avoids the forma-tion of static or slowly moving vapor phase regions or bubbles. Otherwise operation of the process in this mode is similar to that described for ~ig.
hereinabove.
Substantial decreases in the viscosity and pour point of a hydrocarbon feed material and increased API
values are obtained without significant production of coke on the walls of the reaction vessel by practice of the present invention. The following experimental results are provided for the purpose of illustration of the present invention and are not intended to limit the scope of the invention.

.. , __ ... _ , .. _.. _.. . .

~PERIME~TAL I

~ ourteen runs were made to demonstrate direct oxidative heating of a hydrocarbon feed to reduce the 0~ viscosity ~f a Canadian Cold Lake Heavy Oil ~eed. In Run Nos. 1 and 2 the bench-scale simulator described below was used~ For subsequent runs, t~is apparatus was modified as will be explained in detail below. The feed material was held in an oil storaye tank having a 120-volt heater. The feed was fed through a circulat-*

ing pump and a Pulsa-feeder metering pump. The feed material was conducted through three l~-foot tube-in-tube heat exchangers and through a 9-foot tube-in-tube heat exchanger consisting of l/4-inch tubing for the 1~ feed located inside a 1/2-inch tubing $or the product.
The material was then conducted into a fluidized bed sand heater having a 15-inch inner diameter. As the material was introduced into the fluidized bed, the oxidi~ing agent, oxygen, W2S introduced into the feed material line. The material was then conducted through a 5D-foot conduction heating coil section in the fluidized bed and then fed through the 9-foot tube-in-tube heat exchanger and the three 15-foot tube-in-tube heat exchangers. After the thermal exchange, the 2~ material was fed through a series of three ~ressure let-down valves into an expansion separator drum to separate the fluid product from the gaseous product.
In Run No. 3, the system was redesigned so that flow was reversed through the conduction heating coil and the feed entered at the bottom of the coil and exited from the top. Additionally, the oxypen injec-tion apparatus was modified so that oxygen was injected at the bottom of the coil, and a section of l/4-inch tubing was inserted at the oxyyen injection point to 3~ provide a higher velocity for increased mixing.

-2~-* Trademark ~`~ ~

In Run No. 4, the system was modified so that as the oxygen was injected into the feed, the stream flowed through a l-foot section of 3/4-inch tubing.
In Run No. ~ inch Cerefelt aluminum wrap was added to the reactor system as insulation from the 1-0~ foot section of 3/4-inch tubing into the fluidized bed heater.
In Run No. 6, a nitrogen line was added to the system to provide the capability of in;ecting nitrogen instead of oxygen or in combination with oxygen. This run was made with only nitrogen to produce a product sample for comparison with the combustion heating samples.
Run Mos. 7 and 8 used the same apparatus as used in ~un No. 6 with the addition of a second set of check valves and an in-line filter in the oxygen line. These runs started with nitrogen flowing through the system, switching to oxygen when the reaction temperature was reached, and switching back to nitrogen at the end of the run. This procedure allowed for a constant flow of gas to prevent oil from seeping into the oxygen line.
In Run Nos. 9 and 10, the system was modified by introducing the oxygen into the 3/4-inch reactor sec-tion below the introduction point of the feed mate-ial.
Additionally, an in-line filter to the oxygen line was added just below the 3/4-inch reactor section to prevent oil from entering the oxygen line. This ap-paratus was successful in these two runs for preventing oil seepage into the oxygen line.
In ~un No. 11, a l-inch reactor section was sub-stituted for the 3/4-inch reactor section and no oxygen gas was injected into the hydrocarbon feed.
Run No. 12 also used the 1-inch reactor section, and a 7-micron filter frit of sintered stainless steel was used to inject oxygen through the hydrocarbon stream to obtain better oxygen dispersion. This run `3 was ended part way through because the frit became covered with coke material and gas flow into the stream was stopped. Run No. 13 used a 15-micron filter frit.
During this run, a hole was burned in the frit.
In Run No. 14, oxygen was injected through a 1/8-05 inch, 0.049 wall tube and no filter was used.
In Run No. 15, the reactor consisted of 50 feet of 1/4-inch tubing.
Table 1 describes the operating conditions for Run Nos. 1-14 and Table 2 provides a reaction product analysis for Run Nos. 1-14.

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- 1298~03 EXPERIMENTAL II

A product sample from Run No. 5 in Experimental I
was analyzed and compared with oil products obtained by 0!~ indirect heating and with the initial feed material.
The feed material was Canadian Cold Lake Heavy Oil.
The comparison products were identified as Run No. 15 and Feed.
The API gravity and volume percent of various fractions of various materials were compared. Table 3 shows the results of these runs for the feed material, the product from Run No. ~ and Run No. 1~ which was treated by indirect heating.

Table 3 Comparison of Oil Treated by Direct Oxidative Heatinq with Qil Treated by Indirect Heating RUN RUN
FEED NO. 5 NO. 15 20 API Gravity 10.4 12.4 13.2 Vol. % at 430F 1.0 9.9 7.2 Vol. ~ at 430-650F14.322.9 24.9 Vol. % at 650-950F34.233.1 35.0 A mass spectrometric analysis of various oil frac-tions were conducted for the feed material and the products from Run No. 5 and Run No. 15. The results of these tests are shown in Table 4.

3~

Table 4 Direct-Inlet Mass Spectrometric Analysis of Oil Fractions, IBP-430UF Cuts RUN
FEED NO. 5 STRUCTURAL TYPE WT. ~ WT. ~
Paraffins 29.6 34.4 Cycloparaffins 35.1 34.1 Condensed Cycloparaffins 27.5 19.1 Alkyl Benzenes 4.5 9.4 Benzocycloparaffins1.2 1.4 Benzodicycloparaffins0.7 0.6 SUM98.6 99.0 2-Ring Aromatics 1.3 1.
3-Ring Aromatics O.l ---4-Ring Aromatics --- ---5-Ring Aromatics --- ---Other Promatics --- ---Sulfur Condensed Aromatics --- ---Polars ND ND
Not Analyzed - --- ---SUM 1.4 1.0 ND = Not dete~mined.

Table 4 (cont. ?
Direct-Inlet Mass Spectrometric Analysis of Oil ~ractions, 430~-650UF Cuts RUN RUN
FEED NO. 5 NO. 15 STRUCTURAL TYPE WT. ~ WT. ~ WT.
Paraffins 15.7 15.4 16.7 Cycloparaffins 20.5 18.6 15.3 Condensed Cyclo-paraffins 30.9 28.3 24.9 Alkyl Benzenes 9.5 13.1 15.2 Benzocyclo-paraffins 5.7 5.7 7.8 Benzodicyclo-paraffins 4.6 4.6 5.3 SUM86.985.7 85.2 2-Ring Aromatics10.5 10.9 11.3 3-Ring Aromatics1.8 2.0 2.1 4-Ring Aromatics0.1 0.1 0.4 5-Ring Aromatics --- --- ---Other Aromatics --- --- ---Sulfur Condensed Aromatics 0.7 1.2 1.0 Polars ND ND ND
Not Analyzed --- --- ---SUM 13.1 14.2 14.8 ND = Not determined.

~ 129~803 ~

Table 4 (cont.) Direct-Inlet Mass Spectrometric Analysis of Oil Fractions, 650U-950~F Cuts -RUN
FEED NO. 5 STRUCTURAL TYPE WT. % WT. %
Paraffins 11.8 10.8 Cycloparaffins ll.O 10.2 Condensed Cycloparaffins 22.8 22.3 Alkyl Benzenes }2.3 13.8 Benzocycloparaffins 6.7 7.1 Benzodicycloparaffins6.0 6.6 SUM 70.6 70.8 2-Ring Aromatics 17.2 17.8 3-Ring Aromatics 7.2 6.8 4-Ring Aromatics 1.3 1.0 5-Ring Aromatics 0.4 0-3 Other Aromatics --- ---Sulfur Condensed Aromatics 3.3 3.3 Polars ND ND
Not Analyzed ................... --- ---SUM 29.4 29.2 ND = Not determined.

/

~2~38~3 EX~ERIMENTAL II}
The feed material and the product from Run No. 5 were analyzed for the polars content of the 4300F-650F
cuts. The results of this analysis are shown in Table 5. The feed material and the product from Run No. 5 05 were analyzed for concentration of phenols in the 430F-650F fraction. The results of this analysis are shown in Table 6.

Table 5 Polars Contents o 430F-650F Cuts FEED RUN
STRUCTURAL TYPE WT. ~ N0. 5 15 Wt. ~ non polars 67.2 83.5 Wt. ~ non acidic polars 31.114.1 Wt. ~ weak acids1.4 2 0 Wt. ~ strong acids 0.3 0.8 ND = Not Determined ~0 .

.

` ` 129~3803 Table 6 Concentration of Phenols by GC/MS in Weak Acid Fraction Ug/ml (ppm) in Extract RUN
NO. 5 FEED

Methyl phenols 220 180 2-carbon alkyl subst. 480 500 phenols 3-carbon alkyl subst. 1600 560 phenols 4-carbon alkyl subst. 780 940 phenols 5-carbon alkyl subst. 700 360 phenols 6-carbon alkyl subst.' 100 160 phenols Naphthols 170 140 Methyl naphthols 560 300 Dimethyl naphthols 80 ND

-~4-12~03 "

EX~ERIMENTAL IV
~ n elemental analysis of the feed material, the product from Run No. 7, and the product from Run No. 11 was conducted. The results of this analysis are shown in Table 7.
Table 7 Elemental Analysis of Whole Oils, Feed, Run No. 7, and Run No. 11 SAMPLE ELEMENTWT. ~ IN OIL
Feed C 84.04 H 10.42 N 0.50 S 4.65 TOTAL99.61 difference 0.39 H/C ratio1.49 Run No. 7 C 85.00 (oxygen) H 10.22 - N 0.48 ::
S 4.01 TOTAL99.71 difference 0.29 .
H/C ratio1.44 Run No. 11 C &3 9 (indirect H 10.08 -heat) N 0.50 S 4.14 TOTAL98.62 difference 1.38 . . H/C ratio1.44 The feed material, the product from Run Mo. 7, and the product from Run No. 11 were analyzed for sulfur distribution in various fractions of the samples. The results of these analyses are shown in Table 8.

?

Table B
Sulfur Distribution in Oil Samples, Feed Run No. 7, and Run No. 11 -WT.~ S WT.~ S
WT.~ S RUN RUN
DISTILLATION CUT FEED NO. 7 NO. 11 Whole oil 4.65 4.01 4.14 IBP-430F 0.92 2.30 2.34 430-650F 2.47 2.80 3.14 650-950F 3,54 3,90 3 90 950F+ 5.57 5.38 5.50 S in cuts/S in whole 99.4% 96.9% 93.5~ -All values were obtained by X-ray fluroescence.

The feed material, the product from Run No. 7, and the product from Run No. 11 were run through distilla-tions and analyzed with regard to API gravities for various fractions. The results of these runs are shown in Table 9.

Table 9 Distillations and API Gravities of Oils, Feed, Run No. 7, and Run No. 11*
SAMPLE AND CUT API GRAVITY VOL. ~ SUM. VOL.
Feed IBP-430F 32.4 4.5 4.5 430-650F 24.6 13.8 18.3 650-950F 16.3 29.9 48.2 950F+ 3.2 51.8 100.0 Feed contained 1.2 wt. % water; all results on a dry basis.
Feed API gravity was 10.4; IBP was 213F.
Run No. 7 IBP-430F 46.1 14.9 14.9 430-650F 25.0 26.5 41.4 650-950F 13.1 32.4 74.3 950F+ -5.4 25.7 100.0 Feed API gravity was 13.8; IBP was 179F.
Run No. 11 IBP-430F 41.8 25.6 25.6 430-650F 21.7 21.0 46.6 650-950F 12.7 29.8 76.4 950F~ -6.8 23.6 100.0 Feed API gravity was 13.8; IBP was 151F.
*Volume percents were normalized to ïoo~ assumins all losses were in the vacuum residue. In all cases, the material balance was greater than 98~.
Mass spectral structural analyses of the feed material, the product from Run No. 7, and the product from Run No. 11 were conducted for three fractions:
initial boiling point to 430F, 430F to 650F, and 650F to 950F. The results of these runs are shown in Tables 10, 11, and 12.

29~3 ~

Table 10 Mass Spectral Structural Anal~sis of Feed wt percent STRUCTURAL TYPE IBP-430~F 430-650UF 650-950~F
. _ Paraffins 26.7 14.1 9.9 Cycloparaffins28.3 18.1 9.6 Condensed Cyclo-Paraffins 25.3 27.9 18.9 Alkyl Benzenes6.7 9.6 10.1 Benzocyclo-Paraffins 3.8 6.1 6.5 Benzodicyclo-Paraffins 2.2 5 2 5 9 TOTAL 93.0 81.0 60.9 2-Ring Aromatics 5.4 12.9 16.3 3-Ring Aromatics 0.8 3.3 10.0 4-Ring Aromatics --- 1.0 4.9 5-Ring Aromatics --- 0.2 1.0 Other Aromatics 0.3 1.4 3.8 Condensed Aromatic Sulfur Com~ounds 0.3 1.4 3.B
TOTAL 7. O 19 . 0 39 .1 Total Aromatics 19. 7 39 . 9 61.6 Calculated API
Measured API

`~ ? 129~803 -~

Table 11 Mass Spectral Structural Analysis of Run No. 7 wt. percent STRUCTURAL TYPE IBP-43~F 430-650~F 650-950~F

Paraffins 38.5 14.9 10.0 Cycloparaffins33.2 16.4 9.0 Condensed Cycl~-Paraffins 13.3 24.7 17.0 Alkyl Benzenes10.9 13.1 10.5 Benzocyclo-Paraffins 1.8 7.6 6.9 Benzodicyclo-Paraffins 0.8 5.5 5.7 TOTAL 98.5 82.2 59.1 2-Ring Aromatics 1.3 12.4 17.8 3-Ring Aromatics 0.1 2.7 10.7 4-Ring Aromatics 0.1 0.9 4.7 5-Ring Aromatics --- 0.2 2.7 Other Aromatics --- 0.3 0.7 Condensed Aromatic Sulfur Compounds --- 1.3 4.3 TOTAL 1.5 17.8 40.9 Total Aromatics 15.0 44.0 64.0 Calculated API
Measured API

lZg8~03 --Table 12 Mass Spectral Structural Analysis of Run No. 11 wt. percent Paraffins 35.5 13.7 9.7 Cycloparaffins30.7 15.0 8.9 Condensed Cyclo-Paraffins 16.1 2~.1 17.0 Alkyl Benzenes10.5 13.0 10.3 Benzocyclo-Paraffins 3.2 7.1 6.5 Benzodicyclo-Paraffins 1.4 6.2 5.5 TOTAL 97.4 79.1 57.9 2-Ring Aromatics 2.3 14.2 18.1 3-Ring Aromatics 0.3 3.5 11.2 4-Ring Aromatics --- 1.1 4.7 5-Ring Aromatics --- 0.2 2.9 Other Aromatics --- 0.2 0.7 Condensed Aromatic Sulfur Com~ounds --- 1.7 4.5 TOTAL 2.6 20.9 42.1 Total Aromatics 17.7 47.2 64.4 Calculated API
Measured API

EXPERIM~NTAL V
A sample of Canadian Cold Lake heavy oil was processed in a direct oxidative heating pilot simulator. The rëactor consisted of the following three sections: a heat exchanger, a string section, and a reactor section. The heat exchanger was located aboveground and consisted of 240 feet of 1/2-inch tubing inside l-inch tubing. The string section was underground and consisted of 250 feet of 3/8-inch and 1-inch pipe leading _rom ground level down to the reac-tor section. The reactor section was 100 feet long and consisted of 3/8-inch and 3-inch pipe at the bottom of the reactor. All three sections had the smaller .

129~303 ~

diameter tubing concentrically located within the larger diameter tubiny. The hydrocarbon feed flow in the string and reactor sections passed down the inside pipe and returned up the outside pipe.
Sixteen temperature sensing devices were placed at 05 various locations within the reactor. Temperature sen-sor Nos. 1 and 2 were located 100 feet and 200 feet, respectively, down from the ground and monitored the feed temperature. Temperature sensor No. 3 was located near the bGttom of the reactor section, approximately 95 feet from the top of the reactor and measured the product temperature. Temperature sensor Nos. 4 and 5 were located between 95 feet and 7~ feet from the top of the reactor and measured, respectively, the heater temperature and the outside skin temperature of the reactor wall. Temperature sensor No. 6 was located 78 feet from the top of the reactor section and measured the product temperature. Temperature sensor Nos. 8 and 9 were located between 75 feet and 50 feet from the reactor top and measured the product temperature and heater temperature, respectively. Temperature sensor No. 10 was located 50 feet down from the top of the reactor and measured the product temperature. Tempera-ture sensor Nos. 12, 13, and 14 were located less than 50 feet from the top of the reactor section and measured, respectively, the skin temperature, the product temperature, and the heater temperature. Tem-perature sensor Nos. 15 and 16 measured the product temperature and were located 250 feet and 100 feet, respectively, from the surface.
Pressure sensors were also installed in the reac-tor. Pressure sensor No. 1 wzs lo-ated near the bo~tom of the reactor section below the oxidizing agent injec-tion nozzle. Pressure sensor No. 2 was located on the oxidizing agent injection line-p-io- to introduction into the reactor.

. .

303 ~

The in;ector system included li~uid oxygen and nitrogen storage tanks, Sierra flow controllers, a Has-kel air driven compressor, a custom fabricated injec-tion nozzle, and a compressed nitrogen emergency back up system. From the liquid tanks, the gas was passed 05 through evaporators and reyulators set at 175 psi. The gas was then passed through Sierra flow controllers which controlled the flow of each gas to the compres-sor. The capacities of the flow controllers were at 3 scfm for the oxygen line and 6 scfm for the nitrogen line. Separate systems provided for oxygen and nitrogen service to the inlet of the air driven com-pressor. The two gases were combined throughout the remainder of the system. The oxygen and nitrogen were compressed to the system pressure by a ~askell air driven two-stage compressor. The compressor was rated at 5.9 scfm.
The injection nozzle was fabricated by placing a 1/2-inch long plug in the end of a length of 1/4-inch tubing. The plug had previously~been bored with a 1/32-inch diameter hole for the first 1/4-inch and a 1/64-inch diameter hole for the remaining 1/4-inch.
The nozzle was placed vertically pointing upwards half way between the 3-inch outer pipe and the 3/8 inch in-ner pipe. Immediately preceding entry to the 3-inch pipe, a check valve and 5-micron filter were installed to prevent the nozzle from being plugged by foreign particles and to prevent oil from entering the gas line. The nozzle was approximately 25 feet from the bottom of the 98-foot reactor section.
An emergency nitrogen flood system was used to prevent the possibility of hydrocarbon feed ~rom enter-ing the injector line and producing an explosive mix-- ture with subsequent oxygen flow. This back up system consisted of a manifold of six compressed ni~rogen bottles connected to the gas injection line. The ? l2sssn~

compressed nitrogen was isolated from the injection line ~y a solenoid valve connected to a manual switch.
This switch was also connected to-another solenoid valve on the drive air for the Haskell compressor. Ac-tivating this switch caused the compressor to shut down 05 and the compressed nitrogen to flood the injection line.
The reactor section of the system was modified to include an electric heatin~ system. The reactor sec-tion was fitted with 8D0-watt heaters as follows. The bottom section was fitted with 30 bands spaced 3 inches apart, and the top three sections each had 18 bands spaced 14 inches apart.
Throughout the run, the oil feed ~low rate was held nearly constant at 1 gallon per minute and the feed temperature between about 80C and 88C. Canadian Cold Lake Beavy Oil was used as the feed. The system pressure was initially maintained at 1200 psig. During the last half of ~he run, the pressure was gradually reduced to 1900 psig.
The oxygen flow ra~e was 0 for the first 26 hours of the run. It was then stasted at 0.08 scfm, and ove-the next 12 hours, it was gradually increased to 1.2 scfm (37.8 scf/bbl or 3.37 lb/bbl), where it was held for the remainder of the run.
After the initial heating period, the maximum tem-perature was held near 425C for about 1~ hou-s. It was-then raised *o between 43~C and 445C and held thera ~os most of the ne~t 30 hours. The maximum tem-perature was then lowered t~ between 425C and 435~C
3~ for the remainder of the run. Direct o~idati~n of the~
hydrocarbon s~eam provided a ~inal temperatu~e in- -crease of about 2~C to 30C.
Table 13 provides temperature profiles at 1.5 hou-intervals over the run for each of temperature senso_ 3~ NGS. 1-16. Table 14 shows the flow rate of oxvgen and --~3--* Trademar~

? ~ Z~803 ~

nitrogen into the reactor at one and one-half hour in-tervals over the run.

Table 13 Tem~eratures in Direct Oxidative ~eating Pilot Simulator Temperature (C) Time ' 2 3 4 5 6 7 8 9 0:00 --225 309 420 425 401418 422 417 430 1:30 232 311 409 420 393409 413 407 419 3:00 239 321 424 429 405423 427 422 443 4:30 235 318 418 425 401418 422 417 440 6:00 234 317 417 424 399417 421 416 439 7:30 236 319 418 425 401421 ~24 419 - 441 9:00 229 312 416 424 399422 423 421 437 10:30 234 319 420 423 399427 427 426 444 `
12:00 234 318 418 420 396430 430 429 444 13:30 235 316 419 419 396436 428 434 446 15:00 237 327 422 419 3974~0 427 438 459 16:30 237 323 422 418 396440 432 438 461 18:00 236 321 421 417 395439 428 437 459 19:30 234 329 422 417 395441 429 439 461 21:00 236 325 414 407 397443 414 440 446 22:30 238 331 419 412 391443 430 440 452 24:00 238 329 420 413 392448 430 445 458 25:30 240 337 421 412 392451 423 448 459 27:00 226 313 400 403 384408 406 407 441 28:30 245 335 407 407 383436 418 437 454 30:00 2~5 344 416 418 393445 429 444 462 -6~-_ Table 13 (con~.) . _ ~em~era~ures in Direct Oxidative ~eatinq Pilot Simulator . . _ .
Tempera~ure (C) Time 1 2 3 4 5 6 7 8 9 31:30- 245345 -416 417 393 445 426 445 463 33:00 247 349 414~15 392 443 422 441 463 34:30 248 349 413415 391 443 419 442 ~66 36:00 248 356 412414 391 443 420 442 466 37:30 246 355 412415 391 445 418 4~3 466 39:00 247 354 409412 388 444 421 444 464 40:30 246 357 408413 388 443 415 443 463 42:00 246 359 408413 388 443 421 442 464 43:30 248 364 406412 387 442 414 441 463 45:00 218 360 402411 385 439 410 438 461 46:30 249 364 402411 384 433 426 433 458 48:00 249 361 400410 383 436 413 435 459 49:30 249 350 392400 376 435 408 433 452 51:00 246 337 380396 367 424 396 424 441 52:30 242 329 374393 365 A20 392 418 440 54:00 239 325 373396 365 409 397 410 433 55:30 249 343 384407 375 424 412 419 445 57:00 243 336 381405 374 410 412 419 438 58:30 247 344 385409 377 430 405 432 445 60:00 246 352 392415 382 438 416 437 454 61:30 248 347 390413 382 434 416 427 451 63:00 246 346 388413 382 436 408 435 450 64:30 249 352 392416 384 441 414 440 456 66:00 249 353 393416 384 440 413 440 456 67:30 248 348 390415 383 430 422 430 451 69:00 251 361 397420 388 443 417 443 465 70:30 249 354 389407 378 425 414 425 452 72:00 250 358 390411 379 436 4'5 436 4~9 73:30 247 357 396406 381 431 414 437 457 75:00 248 352 393408 380 424 415 423 452 76:30 249 351 393409 381 425 415 426 452 78:00 249 350 394408 381 426 414 427 452 79:30 248 353 395409 382 429 419 429 454 81:00 208 305 381179 167 409 376 426 ~24 82:30 153 227 35083 81 355 301 382 3~2 84:00 112 171 32763 62 307 208 345 348 . . . _ . .

lZ9~8~3 Table 13 (cont.) Temperat~res in Direct Oxidative Heating Pilot Simulator Temperature (C) Time 10 11 12 13 14 15 16 0:00~ 424 434~~~397 414 397 366 252 1:30 410 420 381 395 388 361 260 3:00 427 442 404 419 406 376 262 4:30 419 440 396 410 401 370 260 6:00 420 442 395 407 400 370 260 7:30 421 444 396 409 402 371 262 9:00 417 437 300 404 402 370 257 10:30 426 446 398 412 409 377 259 12:00 424 444 395 410 409 375 259 13:30 424 441 395 408 409 374 259 15:00 429 462 402 421 424 387 259 16:30 433 469 406 426 433 386 258 18:00 432 470 406 427 443 390 260 19:30 436 472 406 428 457 390 259 21:00 425 443 392 410 480 380 262 22:30 423 450 395 411 478 379 262 24:00 432 459 403 424 494 387 265 25:30 429 458 403 425 512 389 265 27:00 411 453 399 415 525 386 260 28:30 418 452 394 416 516 385 275 30:00 426 461 403 424 522 394 275 31:30 427 463 405 426 531 394 274 33:00 425 464 405 426 537 393 276 34:30 426 465 405 426 547 393 280 36:00 426 465 405 426 541 392 278 37:30 430 468 408 428 523 390 278 39:00 427 468 407 427 511 386 278 40:30 429 470 409 429 541 386 280 42:00 426 472 410 430 556 385 280 g3:30 431 472 412 431 560 383 282 45:00 423 469 411 429 563 374 283 46:30 422 468 411 430 568 369 280 48:00 420 469 409 428 573 369 284 49:30 417 459 407 425 465 364 285 51:00 406 448 395 413 408 351 281 52:30 404 448 393 409 401 344 279 54:0~ 397 445 391 409 448 342 277 5~:30 410 452 397 414 469 3~7 285 57:00 404 446 39~ - 4~2 - 472- 347 -28Q
58:30 ~11 450 397 414 ~72 349 283 60:00 419 458 402 419 ~76 354 281 Table 13 (cont.) Temperatures in ~irect Oxidative Heating_Pilot Simulator Temperature (C) . _ _ Time 10 11 12 13 14 15 16 61:30 417 455 403 420 468 355 285 63:00 417 455 403 419 475 355 286 64:30 419 461 406 423 478 357 287 66:00 419 462 406 423 479 357 287 67:30 414 460 406 423 483 357 289 69:00 430 473 413 429 512 362 288 70:30 415 466 411 426 523 359 288 72:00 420 469 410 425 526 359 290 73:30 430 471 416 430 530 365 281 75:00 419 466 410 426 536 359 282 76:30 416 465 410 425 546 359 280 78:00 415 465 410 425 548 359 280 79:30 415 465 410 425 548 359 281 81:00 408 426 409 404 486 311 214 82:30 373 383 376 380 429 234 160 84:00 340 347 343 346 395 187 120 . ~

0~3 Ta~le 14 Flowrates of O~gen and Nitrogen in Direct Oxidative Heatin~ Pilot Simulator Flowrate Flowrate (scfm) (scfm) Time N2 2 TimeN2 2 0:00 1.37 0.01 61:30 0.28 1.19 1:30 1.37 0.01 63:00 0.31 1.19 3:00 1.38 0.00 64:30 0.31 1.19 4:30 1.69 0.01 66:00 0.33 1.19 6:00 1.65 0.01 67:30 0.~9 1.18 7:30 1.60 0.09 69:00 0.31 1.19 9:00 1.45 0.19 70:30 0.31 1.19 10:30 ---- ---- 72:00 0.31 1.19 12:00 1.21 0.~0 ~3:30 0.23 1.19 13:30 1.14 0.69 75:00 0.17 1.20 15:00 0.97 0.79 76:30 0.13 1.20 16:30 0.97 0.79 78:00 0.12 1.19 18:00 0.96 0.79 79:30 0.12 1.19 19:30 0.82 0.89 81:00 1.38 -0.02 21:00 0.45 1.18 82:30 0.36 -0.01 23:15 0.49 1.19 84:00 1.69 -0.02 24:00 0.42 1.19 `
25:30 0.46 1.18 27:00 1.07 0.49 28:30 0.45 1.19 `
30:00 0.42 1.18 31:30 0.33 1.21 33:00 0.33 1.19 34:30 0.33 1.18 36:00 0.36 1.18 37:30 0.32 1.18 39:00 0.34 1.18 40:30 0.34 1.19 42:00 - 0.35 1.18 `-43:30 0.35 1.19 45:00 0.37 1.17 46:30 0.35 1.18 48:00 0.35 1.18 49:30 0.30 1.19 51:00 0.30 1.19 52:30 0.23 1.19 54:00 0.19 1.18 55:30 0.25 1.10 57:00 0.23 1.12 -~
58:30 0.30 1.19 60:00 0.32 1.19 -Table 15 contains data from pressure sensor Nos. 1 and 2 at two ~our intervals over most of the run.

Table 15 Pressures in Direct Oxidative Heating_Pilot Simulator Time Reaction Pressure (psig) 0:00 1330 2:00 1328 4:00 1329 6:00 1333 8:00 1322 10:00 1330 12:00 1324 14:00 1304 16:00 1298 18:00 1305 20:00 1297 22:00 1308 24:00 1308 26:00 1309 --.
28:00 1326 32:00 1303 34:00 1314 36:00 1325 38:00 1350 40:00 1370 :-42:00 1415 44:00 1436 46:00 1491 48:00 1498 :`
50:00 - 1485 52:00 1587 ~ .

~ 803 Eight sample barrels were taken from t~e product stream at r approximately 2S hours, 30 hours, 40 hours, 45 hours, 57 hours, 69 hours, 81 hours, and 92 hours~.
The analytical results of the test run for Barrels 1-8 are provided below in Table 16.
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While various embodiments of the present invention have been described in detail, it is apparent that modifications and adaptations of those embodiments will occur to those skilled in the art. However, it is to be expressly understood that such modifications and 0~ adaptations are within the spirit and scope of the present invention, as set forth in the following claims.

3~

-7~-

Claims (32)

1. A process for reducing the viscosity of a hydrocarbon composition, said process comprising:
a) introducing a feed stream of said hydrocarbon composition into a vessel, said stream comprising a core portion and a boundary layer;
b) increasing the bulk temperature of said stream from a first bulk temperature to a second bulk temperature;
c) introducing an amount of an oxidizing agent into said core portion of said stream to oxidize components in said stream and provide heat to said core portion of said stream to produce a bulk reaction temperature greater than said second temperature;
d) controlling the amount of said oxidizing agent to maintain said reaction tempera-ture below the coking temperature of said feed; and e) maintaining said reaction bulk tempera-ture to produce a reaction product having a lower viscosity than said feed.

2. A process as claimed in Claim 1, wherein said second bulk temperature is at least about 300°C.

3. The method of Claim 1 wherein said reaction temperature is between about 300°C and about 475°C.

4. A process as claimed in Claim 1, wherein said oxidizing agent comprises oxygen.

5. A process as claimed in Claim 1, wherein said hydrocarbon feed is under a pressure above about 1000 psig at said reaction temperature.

6. A process as claimed in Claim 1, wherein the step of increasing the temperature of said stream from the first bulk temperature to the second bulk tempera-ture comprises providing thermal communication between said reaction product and said feed stream.

7. A process as claimed in Claim 1, wherein the differential between said second bulk temperature and said reaction bulk temperature is less than about 35°C.

8. A process as claimed in Claim 7, wherein said differential is less than about 25°C.

9. A process as claimed in Claim 1, wherein the step of introducing an oxidizing agent, comprises in-jecting said oxidizing agent into said stream through an injection nozzle at an injection pressure greater than the pressure of the feed at the point of injec-tion.

10. A process as claimed in Claim 9, wherein said injection pressure is at least about 50 psi greater than the pressure of the feed.

11. A process as claimed in Claim 9, wherein said oxidizing-agent is injected into said stream substan-tially parallel to the line of flow of said stream.

12. A process as claimed in Claim 9, wherein said oxidizing agent is introduced at more than one site in said vessel.

13. A process as claimed in Claim 5, wherein less than about 10 volume percent of said feed stream is in a vapor phase in said reaction zone.

14. A process as claimed in Claim 1, wherein the viscosity of said reaction product is at least about 90 percent lower than the viscosity of said feed.

15. A process as claimed in Claim 1, wherein the API gravity of said reaction product is increased by at least about 2° at 25°C compared to said feed.

16. A process as claimed in Claim 1, wherein the pour point of said reaction product is reduced by at least about 20°C compared to said feed.

17. A method for reducing the viscosity of a hydrocarbon feed by thermal degradation of heavy molecular weight components of the feed at a reaction temperature, said method comprising heating the feed with a heat source to below a reaction temperature and heating the feed to the reaction temperature by inter-nal combustion of a portion of the feed.

18. In a method for reducing viscosity of a hydrocarbon composition using a vertical tube reactor in which an influent stream of hydrocarbon feed is in-creased from a first temperature to a second tempera-ture by heat exchange between said influent stream and an effluent product stream wherein at least one of said streams is in turbulent flow and the pressure on said hydrocarbon feed is increased from a first pressure to a second pressure by the hydrostatic column of said feed the improvement comprising providing an incremen-tal amount of heat to increase the bulk temperature of said feed from said second temperature to a reaction temperature by introducing an oxidizing agent into a core portion of said feed stream to oxidize components in said feed stream.

19. The method of Claim 18 wherein said reaction temperature is between about 300°C and about 475°C.

20. The method of Claim 18 wherein said second bulk temperature is between about 300°C and about 475°C
and said reaction temperature is within about 35°C of said second temperature.

21. The method of Claim 18 wherein said second pressure is at least about 1000 psi.

22. The method of Claim 18 wherein said oxidizing agent is oxygen.

23. The method of Claim 18 wherein said hydrocar-bon composition is selected from the group consisting of whole crude oil, bitumen, kerogen, shale oil and mixtures thereof.

24. The method o Claim 18 wherein said turbulent.
flow is vertical multiphase flow.

25. The method of Claim 18 wherein volatile com-ponents are separated from said effluent product stream and introduced into said influent stream to provide multiphase flow in said influent stream.

26. A method for reducing the viscosity of a whole crude oil said method comprising:
a) passing said oil as an influent stream into the downcomer of a vertical tube reactor to form a column of fluid;
b) bringing said influent stream into heat exchange contact with an effluent product stream both of said streams being in vertical multiphase flow to in-crease the temperature of said influent stream to a heat exchange temperature of between about 300°C and about 475°C;
c) increasing the pressure on said influent stream from an inlet pressure to a reac-tion pressure of at least about 1000 psi by said column fluid;
d) injecting oxygen into the core portion of said influent stream to increase the bulk temperature of said stream to a reaction temperature which is within about 35°C of said heat exchange temperature;
e) maintaining said oil at said reaction temperature to provide a preselected reduction in viscosity of said oil and provide a product; and f) passing said product up a riser as an effluent stream into heat exchange con-tact with said influent stream.

27. The method of Claim l wherein said hydrocarbons are selected from the group consisting of whole crude oil, tar sands oil, bitumen, kerogen, shale oil, and mixtures thereof.

28. The method of Claim 1 wherein the amount of said oxidizing agent is controlled by:
(a) monitoring the bulk temperature of the hydrocarbon stream downstream from an oxidation reaction zone; and (b) adjusting flow of oxidizing agent to maintain said bulk temperature within a preselected temperature range by:
(i) increasing the flow of oxidizing agent when the bulk temperature approaches the lower limit of the preselected temperature range; and (ii) decreasing the flow of oxidizing agent when the bulk temperature approaches the upper limit of the preselected temperature range.

29. The method of Claim 18 wherein the pressure at said reaction temperature is sufficient to maintain the hydrocarbon stream substantially in liquid phase.

30. The method of Claim 29 wherein at least about 90 volume percent of said hydrocarbon stream is in liquid phase.

31. The method of Claim 27 wherein said hydrocarbon feed stream contains less than about 13 weight percent water.

32. A vertical tube reactor apparatus suitable for reducing the visoosity of hydrocarbons which comprises:
(a) a downcomer for receiving an influent stream of hydrocarbons and bringing said stream into heat exchange contact with an effluent product stream and provide a heated influent stream;
(b) a reaction zone in fluid communication with said downcomer to receive said heated influent stream;
(c) an oxidant injection nozzle located within said reaction zone and oriented to introduce oxidant into said heated influent stream substantially parallel to the flow of said stream and increase the temperature of the stream-