CATALYTIC DOWN-HOLE UPGRADING OF HEAVY OIL AND OIL SAND BITUMENS

Abstract

The invention relates to systems and methods for catalytic down-hole upgrading of heavy oil and oil sand bitumens. The method enables upgrading heavy oil in a production well within a hydroprocessing zone including the steps of: introducing a controlled amount of heat to the hydroprocessing zone; introducing a selected quantity of hydrogen to the hydroprocessing zone to promote a desired hydrocarbon upgrading reaction; and, recovering upgraded hydrocarbons at the surface. The invention further includes the hardware capable of performing the method.

Description

FIELD OF THE INVENTION

The invention relates to systems and methods for catalytic down-hole upgrading of heavy oil and oil sand bitumens. The method enables upgrading heavy oil in a production well within a hydroprocessing zone including the steps of: introducing a controlled amount of heat to the hydroprocessing zone; introducing a selected quantity of hydrogen to the hydroprocessing zone to promote a desired hydrocarbon upgrading reaction; and, recovering upgraded hydrocarbons at the surface. The invention further includes the hardware capable of performing the method.

BACKGROUND OF THE INVENTION

Heavy oil and bitumen are highly viscous oils that are difficult to produce and require upgrading before being sent into refineries. Costly processing methods and problematic transportation make heavy oils a reasonable substitute for conventional oil only when energy prices are high enough to justify their processing costs. Due to the sharp increase in oil prices in the 21st century, special attention has been paid to heavy oil and bitumen resources.

Heavy oil like any kind of petroleum or crude oil consists of a wide range of constituents mainly mixtures of hydrocarbons and compounds containing sulfur, nitrogen, oxygen and metals. Metals are typically vanadium and nickel and their percentage in the petroleum increases in most viscous oils. The physical properties of petroleum vary widely depending on its constituents and their amounts.

The definition of a heavy oil is normally based on its API gravity and/or viscosity and is quite arbitrary. A commonly accepted definition classifies the heavy oil as petroleum whose API gravity is between 10 and 20. Also, heavy oils usually and not always have sulfur contents of higher than 2 wt %. In the case of bitumen, the term refers to the oil produced from bituminous sand formations that are used to recover the bituminous material by mining operations.

The physical properties and chemical composition of oil change not only based on geographical location but also with the depth of the oil in a particular location and will normally contains small amounts of organic materials containing sulfur, oxygen and nitrogen as well as metallic compounds such as nickel, iron, and copper. For lighter oils the amount of hydrocarbon molecules can be as high as 97% and for heavy oil and bitumen can be as low as 18%. Table 1 shows a range of each main constituent of a typical light and heavy petroleum:

TABLE 1
Light and Heavy Petroleum Constituents
	Element	Light	Heavy
	Carbon:	83.0% to 87.0%	83.4 ± 0.5%
	Hydrogen:	10.0% to 14.0%	10.4 ± 0.2%
	Nitrogen:	0.1% to 2.0%	0.4 ± 0.2%
	Oxygen:	0.05% to 1.5%	1.0 ± 0.2%
	Sulfur:	0.05% to 6.0%	5.0 ± 0.5%
	Metals (Ni and V)	<1000 ppm	>1000 ppm

Oil is also classified in the following general classifications based on the average boiling point of its constituents.

Gases & Naphtha: The main constituent of petroleum gases is methane. The other major hydrocarbons are ethane, propane, butane, isobutene and some C4+ alkanes.

Middle Distillates: The main portion of this fraction is comprised of the saturated species; however, aromatics and heterocyclic compounds represent a considerable portion. Most of the aromatics are di- and tri-methyl naphthalenes. The percentage of the sulfur molecules is very low. Trace amounts of both basic and neutral nitrogen compounds are normally present. The boiling range is typically between 180° C. and 340° C.

Vacuum Gas Oils: The boiling range for the vacuum gas oil (VGO) is between 340° C. and 540° C. The quantity of the aromatics (mono- and di-aromatics) is greater than the saturates. Saturated compounds in VGO consist of iso-paraffins and naphthene species. Heterocyclic compounds are also significant in VGOs. The major sulfur compounds are the thiophenes (mostly benzothiophenes and dibenzothiophenes) and thiacyclane sulfur which are present in greater amounts than sulfide sulfur.

Vacuum Residue: This fraction is the most complex fraction and contains the majority of heteroatom molecules of the petroleum whose boiling points are over 540° C. Characterization of individual constituents of this fraction may utilize various empirical/semi-empirical thermo-physical models based on the observed functionalities, apparent molecular weight and elemental analysis.

Heavy Oil Production from Oil Sands

Oil sands are deposits of bitumen and are found in 70 countries with three quarters of world reserves located in Canada and Venezuela. Production of heavy and extra-heavy oils in Canada started some 38 years ago by surface mining the Athabasca oil sands. As of Dec. 31, 2005, Canada's proven oil reserves are estimated as 177.9 billion barrels from which 173.7 billion is in the form of oil sands.

The production of bitumen from oil sands has been revolutionized in the recent years. Today less than 10% of bitumen production occurs through mining. Enhanced oil recovery techniques (discussed below) enable producing reserves that are deeper than 75 meters under the ground surface. Most of the in-situ production of bitumen and heavy oil comes from deposits buried more than 400 meters.

The most common in-situ production methods are Cyclic Steam Stimulation (CSS) and Steam Assisted Gravity Drainage (SAGD) which require the injection of hot fluids into the reservoir. Canada's largest in-situ bitumen production projects are at Cold Lake where deposits are heated by steam injection and the bitumen is brought to the surface and later diluted with condensates for pipeline transportation.

In-situ technology in recent years has been a subject of study for use in down-hole upgrading processes. The in-situ upgrading that has been targeted in research projects includes hydrotreating of oil, mostly hydrodenitrogenation and hydrodesulfurization, hydrocracking and asphaltene precipitation. These upgrading projects mostly do not provide a high degree of upgrading in terms of contaminant removal and API gravity increase.

Traditionally oil recovery methods are divided into three categories: primary, secondary and tertiary based on their chronological order in production. Primary recovery is accomplished by the natural energy of the oil reservoir, secondary recovery is based on the energy of injection of gas or water to displace oil towards the production well and tertiary recovery use miscible gases, chemicals and thermal energy to displace the oil and are implemented when secondary methods become uneconomical.

As heavy oil reservoirs cannot be produced based on their natural energy or waterflooding, the term EOR (Enhanced Oil Recovery) is used to describe chemical and thermal methods for heavy oil recovery. Examples of EOR processes are injections of gases such as nitrogen and CO₂, injection of liquid chemicals including polymers, surfactants and hydrocarbon solvents and thermal methods such as steam or hot water injection or in-situ generation of thermal energy.

Upgrading

Upgrading is generally defined as any treatment of bitumen or heavy oil that increases its value. Therefore, the minimum objective is to reduce the viscosity of oil and the maximum objective is to obtain a crude oil substitute of higher quality.

Hydroprocessing reactions are thermal processes that take place in the presence of hydrogen. Such reactions can be both destructive and non-destructive. Destructive hydrogenation (hydrogenolysis and hydrocracking) is conversion of higher molecular weight compounds to lower-boiling point compounds. Non-destructive hydrogenation or hydrotreating are simple hydrogenation reactions during which the quality of oil improves by removing certain contaminants of oil from its molecular structure such as sulfur (hydrodesulfurization (HDS)), nitrogen (hydrodenitrogenation (HDN)) and metals. Generally speaking the reactivity of the hydroprocessing reactions increases in the following order:

Hydrocracking<HDN<HDS

The reaction conditions vary in the different processes, however a typical temperature range is 300-345° C. and the hydrogen partial pressure can be in the range of 500 to 1000 psi.

The catalyst used in hydrotreating reactions is normally cobalt-molybdenum with typically 10% molybdenum oxide and less than 1% cobalt oxide and the support is alumina. However, a wide range of metals can be effective: cobalt, iron, nickel promoted copper and copper chromite. The type of catalyst that is used in each process can change based on the immediate objective. For example, CoMo type formulae are generally used for HDS reactions, the NiMo type are employed for hydrogenation and HDN reaction and the NiW type are used for hydrogenation of very low sulfur cuts. The existence of sulfur and metals is a challenge in hydrotreating reactions because these compounds poison the catalyst. Therefore, more resistant catalysts are used such as CO—Mo—Al₂O₃.

Hydrocracking is the reaction between hydrogen and oil fractions, mostly vacuum distillates and residue, in which the reactants crack into lighter fractions. Based on the objective chosen with respect to the extent of conversion and the quality of the products, hydrocracking can be divided into mild hydrocracking and conventional hydrocracking. Both of these two hydrocracking processes are similar with respect to the reactions, however, the products and their quality can vary because of the different reaction conditions. Mild hydrocracking normally takes place at some 50-80 bar (5-8 MPa) total pressure and temperature of 350-430° C., where conventional hydrocracking, the total pressure is about 100-200 bar (10-20 MPa) and the temperature is between 380-440° C. The mechanism of hydrocracking is similar to that of catalytic cracking but includes concurrent hydrogenation. The products of hydrocracking are either saturated or aromatic rings but not olefin.

An important hydrocracking reaction is the partial hydrogenation of polycyclic aromatics and the ultimate rupture of saturated rings to monocyclic aromatics. In the case of residue, hydrocracking is used for processes such as desulfurization and residue conversion to lower boiling distillates.

The reactions take place in the presence of dual-function catalysts. Silica-alumina catalyst promotes cracking reactions where platinum, tungsten oxide, or nickel contributes to hydrogenation reactions.

One major difference between hydrocracking and hydrotreating is the residence time and the extent of the decomposition of the non-heteroatom constituents. The upper limits of hydrotreating often overlap with the lower limits of hydrocracking.

The main advantages of down-hole upgrading include the reduction in refinery and upgrading costs, the reduction in size of surface upgrading facilities and the utilization of the pre-introduced heat from thermal processes.

Examples of past systems and methods of upgrading oil include U.S. Pat. No. 6,964,300 which describes a process of in situ thermal recovery from a permeable formation using a heater within the wellbore; U.S. Pat. No. 6,742,593 which describes in situ thermal processing of a hydrocarbon containing formation using heat; U.S. Pat. No. 7,121,342 which describes thermal processes for subsurface formations; U.S. Pat. No. 6,996,374 which teaches a method of increasing hydrocarbon mobility within a permeable formation using gas; US patent application 2003/0062163 which teaches a method of in situ upgrading; U.S. Pat. No. 3,817,332 which teaches a method and apparatus of catalytically heating a wellbore; US Patent application 2006/017053 which teaches a process for improving extraction of crude oil by circulating hot hydrocarbons in order to heat an underground reservoir; U.S. Pat. No. 6,056,050 which teaches injecting steam through a horizontal well into a formation for enhancing viscous oil recovery; US Patent application 2006/0231455 which teaches a method for producing and upgrading oil; US Patent application 2005/0239661 which discloses a downhole catalytic combustor for enhanced heavy oil mobility; and U.S. Pat. No. 4,597,441 which teaches a recovery of oil by in situ hydrogenation.

SUMMARY OF THE INVENTION

In accordance with the invention, there is provided a method of upgrading heavy oil in a production well within a hydroprocessing zone comprising the steps of: introducing a controlled amount of heat to the hydroprocessing zone; introducing a selected quantity of hydrogen to the hydroprocessing zone to promote a desired hydrocarbon upgrading reaction; and, recovering upgraded hydrocarbons at the surface.

In further embodiments of the invention, a catalyst is introduced to the hydroprocessing zone where the catalyst may be a nano-particle catalyst that may be circulated within the hydroprocessing zone.

The hydroprocessing zone is preferably a vertical section of a wellbore where heavy oil is preferably separated from water prior to introducing heavy oil into the hydroprocessing zone. In various embodiments, the catalyst is a bi-metallic catalyst of the general formula: B_xM_yS_{[(1.1 to 4.6)y+(0.5 to 4)x]} where B is a group VIIIB non-noble metal and M is a group VI B metal and 0.05≦y/x≦15. In one embodiment, 0.2≦y/x≦6 or y/x=3. The catalyst may be tri-metallic catalysts of the general formula: B_xM1_yM2_zO(2 to 3)_zS_{[(0.3 to 2)y+(0.5 to 4)x]} where B is a group VIIIB non-noble metal and M1 and M2 are group VI B metals and 0.05≦y/x≦15 and 1≦z/x≦14. In one embodiment the y/x ratio is in the range of 0.2<y/x<6 and in more specific embodiments 10<z/x<14 or z/x=12. In another embodiment, 1<z/x<5 and the upgrading process is mild hydrocracking. In another embodiment, z/x=3 and the upgrading process is mild hydrocracking. Other upgrading reactions may be hydrodenitrogenation and hydrodesulfurization.

Heat may be introduced to the one or more hydroprocessing zones (where different hydroprocessing reactions may occur) using any one of or a combination of electrical, hot fluid, or an in-well combustion device.

In another embodiment, the method of the invention, controls the reaction parameters in different areas of the hydroprocessing zone so as promote different hydroprocessing reactions in different areas.

In yet another embodiment, the upgrading process is part of a steam flooding process including any one of steam assisted gravity drainage (SAGD), vapor extraction (VAPEX), cyclic steam stimulation (CSS) and CAPRI.

In another embodiment, the invention provides a system for upgrading heavy oil in a production well within a hydroprocessing zone comprising: a downhole heater for introducing a controlled amount of heat to the hydroprocessing zone; a hydrogen delivery system for introducing a selected quantity of hydrogen to the hydroprocessing zone to promote a desired hydrocarbon upgrading reaction; and, a surface recovery system for recovering upgraded hydrocarbons at the surface. The system may also include a downhole water separator for separating water from heavy hydrocarbon, the downhole water separator operatively located upstream of the hydroprocessing zone.

DESCRIPTION OF THE DRAWINGS

The invention is described with reference to the drawings in which:

FIG. 1 is a schematic diagram of an in situ upgrading system having horizontal and vertical wells;

FIG. 2 is a schematic diagram of an in situ upgrading system where production and injection take place through the same well;

FIG. 3 is a schematic diagram of an in situ upgrading system where production and injection wells are located a distance from each other;

FIG. 4 is a schematic diagram of an in situ upgrading system using a THAI configuration;

FIG. 5 is a schematic diagram of a two-stage separator within a wellbore;

FIG. 6 is a schematic diagram of horizontal wellbore segments and feed streams entering each segment;

FIG. 7 is a diagram of a HYSYS interface showing horizontal wellbore segments and entering feed streams;

FIG. 8 is a schematic diagram of vertical wellbore segments in series configuration;

FIG. 9 is a diagram of a HYSYS interface showing the vertical wellbore and the inlet streams;

FIG. 10 is diagram showing a network of hydrocracking reactions between various oil fractions;

FIG. 11 is a diagram showing a simplified hydrocracking network;

FIG. 12 shows HDS conversion percent for non-residue lumped fractions (Diam. 15 cm—production rate 1.39 m3/h);

FIG. 13 shows HDS conversion percent for residue fraction (Diam. 15 cm—production rate 1.39 m3/h);

FIG. 14 shows HDN conversion percent for non-residue lumped fractions (Diam. 15 cm—production rate 1.39 m3/h);

FIG. 15 shows HDN conversion percent for residue fraction (Diam. 15 cm—production rate 1.39 m3/h);

FIG. 16 shows HDS and HDN conversion percent at 350° C. for non-residue lumped fractions (Diam. 15 cm—production rate 1.39 m3/h);

FIG. 17 shows HDS and HDN conversion percent at 350° C. for residue fraction (Diam. 15 cm—production rate 1.39 m3/h);

FIG. 18 shows weight percent of residue sulfur compounds, non-residue lumped fractions sulfur compounds and the total sulfur compounds in feed and HDS product streams at various wellbore lengths at 350° C. (Diam. 15 cm—production rate 1.39 m3/h);

FIG. 19 shows weight percent of residue nitrogen compounds, non-residue lumped fractions nitrogen compounds and the total nitrogen compounds in feed and HDN product streams at various wellbore lengths at 350° C. (Diam. 15 cm—production rate 1.39 m3/h);

FIG. 20 shows Composition of feed (typical Alberta bitumen);

FIG. 21 shows Volume percent change due to hydrocracking on conventional catalyst at 425° C. (Diam. 15 cm) —SOR 0;

FIG. 22 shows Volume percent change due to hydrocracking on conventional catalyst at 350° C. (Diam. 15 cm) —SOR 0;

FIG. 23 shows volume percent change due to hydrocracking on conventional catalyst at 375° C. (Diam. 15 cm) —SOR 0 ;

FIG. 24 shows volume percent change due to hydrocracking on conventional catalyst at 403° C. (Diam. 15 cm) —SOR 0;

FIG. 25 shows volume percent change due to hydrocracking on conventional catalyst at 100 m wellbore (Diam. 15 cm) —SOR 0;

FIG. 26 shows volume percent change due to hydrocracking on conventional catalyst at 200 m wellbore (Diam. 15 cm) —SOR 0;

FIG. 27 shows volume percent change due to hydrocracking on conventional catalyst at 300 m wellbore (Diam. 15 cm) —SOR 0;

FIG. 28 shows Volume percent change due to hydrocracking on conventional catalyst at 500 m wellbore (Diam. 15 cm) —SOR 0;

FIG. 29 shows Volume percent change due to hydrocracking on conventional catalyst at 425° C. (Diam. 10 cm) —SOR 0;

FIG. 30 shows Volume percent change due to hydrocracking on conventional catalyst at 350° C. (Diam. 10 cm) —SOR 0;

FIG. 31 shows Volume percent change due to hydrocracking on conventional catalyst at 375° C. (Diam. 10 cm) —SOR 0;

FIG. 32 shows Volume percent change due to hydrocracking on conventional catalyst at 403° C. (Diam. 10 cm) —SOR 0;

FIG. 33 shows Volume percent change due to hydrocracking on conventional catalyst at 100 m wellbore (Diam. 10 cm) —SOR 0;

FIG. 34 shows Volume percent change due to hydrocracking on conventional catalyst at 200 m wellbore (diam. 10 cm) —SOR 0;

FIG. 35 shows Volume percent change due to hydrocracking on conventional catalyst at 300 m wellbore (Diam. 10 cm) —SOR 0;

FIG. 36 shows Volume percent change due to hydrocracking on conventional catalyst at 500 m wellbore (Diam. 10 cm) —SOR 0;

FIG. 37 shows Volume percent change due to hydrocracking on conventional catalyst at 350° C. (Diam. 15 cm) —SOR 1;

FIG. 38 shows Volume percent change due to hydrocracking on conventional catalyst at 375° C. (Diam. 15 cm) —SOR 1;

FIG. 39: Volume percent change due to hydrocracking on conventional catalyst at 403° C. (Diam. 15 cm) —SOR 1;

FIG. 40 shows Volume percent change due to hydrocracking on conventional catalyst at 425° C. (Diam. 15 cm) —SOR 1;

FIG. 41 shows Volume percent change due to hydrocracking on conventional catalyst at 100 m wellbore (Diam. 15 cm) —SOR 1;

FIG. 42 shows Volume percent change due to hydrocracking on conventional catalyst at 200 m wellbore (Diam. 15 cm) —SOR 1;

FIG. 43 shows Volume percent change due to hydrocracking on conventional catalyst at 300 m wellbore (Diam. 15 cm) —SOR 1;

FIG. 44 shows Volume percent change due to hydrocracking on conventional catalyst at 500 m wellbore (Diam. 15 cm) —SOR 1;

FIG. 45 shows Volume percent change due to hydrocracking on conventional catalyst at 350° C. wellbore (Diam. 15 cm) —SOR 10;

FIG. 46 shows Volume percent change due to hydrocracking on conventional catalyst at 375° C. wellbore (Diam. 15 cm) —SOR 10;

FIG. 47 shows Volume percent change due to hydrocracking on conventional catalyst at 403° C. wellbore (Diam. 15 cm) —SOR 10;

FIG. 48 shows Volume percent change due to hydrocracking on conventional catalyst at 425° C. wellbore (Diam. 15 cm) —SOR 10;

FIG. 49 shows API gravity increase for SOR 0—conventional catalyst;

FIG. 50 shows API gravity increase for SOR 1—conventional catalyst;

FIG. 51 shows API gravity increase for SOR 10—conventional catalyst;

FIG. 52 compares the API gravity increase at 425° C. for SOR 0 and SOR 1—conventional catalyst;

FIG. 53 compares the API gravity increase at 403° C. for SOR 0 and SOR 1—conventional catalyst;

FIG. 54 shows volume percent change due to hydrocracking on UD catalyst at 425° C. wellbore (Diam. 15 cm)

FIG. 55 shows Volume percent change due to hydrocracking on UD catalyst at 350° C. wellbore (Diam. 15 cm);

FIG. 56 shows Volume percent change due to hydrocracking on UD catalyst at 375° C. wellbore (Diam. 15 cm);

FIG. 57 shows Volume percent change due to hydrocracking on UD catalyst at 403° C. wellbore (Diam. 15 cm);

FIG. 58 shows API gravity increase for SOR 0—UD catalyst;

FIG. 59 compares the API gravity increase at 425° C. for SOR 0

FIG. 60 compares the API gravity increase at 403° C. for SOR 0; and,

FIG. 61 compares the API gravity increase at 375° C. for SOR 0.

DETAILED DESCRIPTION OF THE INVENTION

In accordance with the invention and with reference to the figures, systems and methods for upgrading hydrocarbons within a petroleum reservoir are described. In particular, the methods enable upgrading of heavy, extra heavy and shale oils and bitumen within a production well bore using selective downhole heating elements, hydrogen and catalyst injection so as to integrate exploitation with in-situ upgrading. The methods of the invention are particularly applicable to steam-assisted gravity drainage (SAGD) and vapor extraction (VAPEX), and cyclic steam stimulation (CSS) recovery methodologies.

In a preferred embodiment, as shown in FIG. 1, the invention provides a system for hydrocarbon upgrading in a well bore system having both horizontal and vertical sections. As discussed below, the methodologies of the invention may be applied to other EOR techniques including wells having only a single vertical section. As shown in FIG. 1, the horizontal section 10 serves to collect the hot oil/water mixture feed 11 via perforations 12 on its surface, with any one of or both of the horizontal section 10 and/or vertical section 14 serving as a reactor with reactor elements. The temperature of the feed is increased by heat introduced to the body of the well by any one of or a combination of electrical, combustion, hot gases or other localized heaters 15 in accordance with various EOR techniques discussed below. Within the horizontal and/or the vertical section, the feed 11 may be mixed with hot hydrogen injected into the well via a gas liner 16 within the vertical section 14 and/or horizontal sections 10 of the well.

In Situ Upgrading

In accordance with the invention and with reference to the figures, down-hole or in-situ upgrading processes are described for various EOR methodologies. Such processes have been a previously unsuccessful alternative to conventional upgrading processes as a result of the difficulties of placing catalyst underground, treating the abundant amounts of brine, high partial pressure of steam and low partial pressure of hydrogen. However, as detailed below, there are advantages in down-hole upgrading by employing the down-hole energy (up to 35 MPa and 80° C.), and the porous media (mineral formation) that can act as a natural chemical catalytic reactor.

Generally, the process of in-situ upgrading includes:

• • Placement or formation of catalyst in an oil bearing medium;
  • Mobilization of oil components over the catalyst;
  • Introducing co-reactants such as hydrogen to the reaction environment; and,
  • Creating the necessary conditions for the reactions;

There are two main routes for chemical reaction upgrading, namely the addition of hydrogen, or hydrogen donors and carbon rejection. Hydrogen addition results in hydrogenation and little carbon deposition because the by-products are mostly hydrogen sulfide or light hydrocarbons which are mostly gaseous that will automatically exit the wellbore. Gases may also contribute to production increases because of their miscibility in the oil and contribution to reducing viscosity. Another advantage of this process is hydrogenation of carbon deposits.

Carbon rejection is beneficial as it leaves highly carbonaceous materials in the wellbore and produces upgraded oil. However, the deposition of such materials in the reaction medium will contribute to wellbore plugging and catalyst deactivation. Furthermore, most oil reservoirs also contain significant amounts of brine that will have a significant effect in down-hole upgrading of heavy oils.

Further still, any injected gas, in order to have a reasonable partial pressure to react, must have a higher pressure than that of injected steam. As the saturated steam pressure at 300° C. is about 1235 psi, for a wellbore with reaction temperatures of 250-350° C. to contain steam, a depth of 1150 to 2800 feet is required. In other words at such conditions the gas must be injected at over 1200 psi for a length of some 2000 feet.

Different processing scenarios can be implemented for down-hole upgrading:

In one embodiment, a catalyst bed is placed in an oil-bearing interval by gravel packing, proppant injection, or water injection as shown in FIG. 2. Oil flows over the catalyst either naturally or by induced drive mechanisms. Oil is produced through the perforations in the well casing and is directed to the surface by the production tubular. An injection tubular is used to inject heated fluids, such as hydrogen or hydrogen donors into a volume below the catalyst bed.

In another embodiment, as shown in FIG. 3, the catalyst is placed in close vicinity to the production well; however, the injection process is through the injection well, placed a further distance from the production well. Thermal drive is induced by a combustion front. The combustion produces hydrogen and carbon monoxide which mobilizes the oil front. At the same time, water transfers heat ahead of the combustion front by steam override. In this configuration, additional heat can also be provided through the production zone as per FIG. 1.

In a further embodiment as shown in FIG. 4, catalyst is placed inside or around a horizontal production well, where a vertical injection well injects hot air into the reservoir. This method is called CAPRI and is the catalytic form of Toe to Heel Air Injection (THAI) method. The upgrading extent is higher than THAI because of the use of the catalyst in the system.

In a further embodiment, Steam Assisted Gravity Drainage (SAGD) (not shown) methodologies are utilized with catalyst placed or injected inside the wellbore. In a SAGD production method, two parallel wells exist including an upper well for steam injection well and a lower one as the production well. The steam injection well drives the oil into the other horizontal well underneath. Oil contacts catalyst placed inside the production well, and the upgrading reactions are promoted as the oil moves towards the vertical well. Catalyst can either be a solid fixed bed or an ultra dispersed catalyst in the liquid phase.

Brine

As previously mentioned, the presence of significant volumes of brine poses a major challenge in down-hole upgrading. DOWS (down-hole oil water separation) technology typically consists of a separation and a pumping stage. In one design (not shown), gas is separated from the liquid through gravity separation without introduction of any centrifugal force, nozzles or other types of mechanisms utilizing the difference in the density of the two phases as the factor in the separation. This design involves the installation of a pump intake below the lowest point of fluid entry into the wellbore and requires an open casing-tubing annulus along the wellbore. The gas bubbles rise through the liquid phase and leave its surface and move upwards in the casing-tubing annulus. The liquid phase is accumulated at the bottom of the well and enters the pump intake to be discharged into the tubing.

In an alternative design, separation occurs in two stages as shown in FIG. 5. In the first stage, separation of gas from liquid occurs in the wellbore tubing-casing annulus. The gas bubbles leave the gaseous liquid in the annulus and move upwards in the casing. The remaining mixture of gas and liquid enters the second stage down-hole gas separator through an anchor port and its perforations on its surface leading to further separation. In this system, the amount of gas that flows with the liquid into the tube and to the pump intake is minimized.

Upgrading Conditions

Pressure

Hydrostatic pressure is the means to obtain the upgrading reaction pressures and is calculated from

P=ρgh

where ρ is the density of the ground above, g is the earth's gravity and h is the depth of the wellbore.

Energy

To provide the energy to the reaction medium various methods can be utilized such as introducing hot fluids, steam injection or fire flooding. Other methods use point sources including down-hole steam generation or combustion, electromagnetic stimulation and down-hole heating with electric coils.

Down-hole gasification or combustion may be utilized for sub-terrain heating. As noted a mixture of fuel and air are injected into the wellbore and are ignited creating a front that moves towards the production well. Wet oxidation can been used to inject steam under a formation at 315-340° C. and 2000-2500 psi. A heat conductive system that employs a down-hole gas-fired burner is capable of heating a transfer fluid to 815-1400° C. Another benefit of down-hole heating is the generation of CO which in proper conditions controls the extent of oxidation (together with combustion or partial oxidation catalysts), produces H₂ through the water-gas shift reaction.

A particular benefit is that operation of a down-hole steam generator or gasifier below a catalyst zone produces upward flow of heat and combustion gases that can provide heat and H₂ (or CO) for upgrading.

Catalyst

Catalyst is normally placed in the vicinity of the well either by injecting a liquid phase solution or adding solid catalyst particles around the wellbore.

In the case of solid particles, and where recovering the catalyst is not practical, used or regenerated hydroprocessing catalysts may be appropriate for placement. The major problem associated with solid-phase catalyst is the collection of impurities on the catalyst area resulting in the deactivation of catalyst and also greater pressure drops with time.

Injection of homogenous catalyst occurs in the area surrounding the production well. Fluid phase catalyst reaches further distances from the wellbore; therefore, an advantage is that if plugging occurs due to the reactions, it will be away from the wellbore and will have less severe effects on production rates.

Homogenous catalysts mostly have similar active metals to those of heterogeneous catalysts, mainly molybdenum and iron. The additives are mostly cobalt and nickel and the sulfide metal is the active phase. An advantage to using a fluid phase catalyst is that it can be prepared in remote areas. For example amines such as ethylene diamine can be added to aqueous ammonium heltamolybdate and cobalt nitrate mixtures which stabilizes the solution and allows the metals to be deposited in areas remote from solution preparation.

Homogenously dispersed catalysts can also be used for combustion catalytic upgrading. Aqueous phase iron or tin salts dispersed in a mixture of sand/oil/water in a combustion tube experiment resulted in increased fuel deposition, higher velocity of combustion front and lower oxygen combustion.

Methodology

A wellbore consisting of both horizontal and vertical sections was studied as per FIG. 1 although it is understood that the upgrading technologies described herein may be applied to other EOR techniques as understood by those skilled in the art. The horizontal well collects the mixture of oil and steam via the perforations on its surface and directs them to the vertical section. The total length of the horizontal section may be varied based on the well location and reservoir length. In this description, the length of this section was assumed to be 1000 meters. The vertical well may also have different lengths which will result in various residence times for the upgrading reactions. The down-hole temperature of oil was assumed to be 220° C. as a typical high temperature of SAGD steam injection.

To evaluate the effectiveness of in-situ upgrading, a HYSYS simulator was used (Aspen Tech., Houston). This simulator offers a comprehensive Oil Manager which allows introducing various oil assays to the model and the creation of pseudo-components based on the desired assay. Also the reaction package provides a high level of control over the reaction stoichiometry, kinetics, units and the phases. Finally the simulator encompasses a comprehensive set of objects and unit operations that permit simulation of the wellbore with high level of control over each segment.

As shown in FIG. 1, the model consists of the vertical and horizontal sections of a wellbore and the upgrading sections. The oil enters the wellbore through the perforations and the entrance point. The fluids then pass through an optional steam/water separator 50, followed by flow into the vertical wellbore where the oil is mixed with the injected stream of ultra dispersed catalyst and hydrogen gas. The mixture of oil, catalyst and hydrogen moves upwards where it is heated and the hydroprocessing reactions take place. At the end of the vertical well, the produced oil is partially upgraded.

HYSYS Fluid Package

Simulation models in HYSYS are created based on previously defined fluid packages. The choice of any package is based on the specific system under consideration (the components that are involved and their interaction) and also the operating conditions.

The main thermodymic package choice is either the Equation of State or the Activity Model. The Equation of State chosen for this model is Peng-Robinson which was developed originally to deal with hydrocarbon gas models. This model has been shown to be very efficient for most hydrocarbon based fluids over a wide range of operating conditions.

Simulation Components

In a flowsheet simulation, components can either be defined as pure or pseudo-components. Pure components are specific chemical compounds such as water or hydrogen. The pseudo-components, which are called hypothetical components in HYSYS, are not pure but are treated as those. Their definition is based on the objective and nature of the simulation and can vary.

In refinery simulations, a major advantage of defining pseudo-components is to limit the number of components in the system by grouping them into limited number of groups. This significantly decreases the computation time needed for analyzing the components of a stream without affecting the accuracy.

The components are defined based on their properties which are usually required for thermodynamic calculations. These required properties vary in different simulators but some common ones are the critical temperature, pressure and volume, acentric factor, solubility parameter, liquid molar volume, van der Waals area and volume and latent heat of vaporization. There is no need to input these properties for the pure components because they exist in the simulator's database; however, the case is different for pseudo-components and their properties are normally estimated by the correlations and some major input properties, usually normal boiling point, specific gravity and molecular weight.

Oil Components

The Oil Manager in HYSYS is used to input the characteristics of the oil assay in general and the feed in particular. The main input data is the True Boiling Point (TBP) distillation curve which is obtained as part of the assay. This requires inputting the boiling points of each fraction and the corresponding volume percents in the liquid form at a specific pressure.

Once the assay is defined, the ‘Blend’ tab describes the feed. Any blend in this section will be an arbitrary mixture of oil fractions. The blend is defined as follows:

A temperature is input and the number of cuts with End Boiling Points of less than this value is input. For example to define the naphtha cuts, the cut End Point of 204° C. is input and then in the cell across the row, number 1 is input showing that there is one cut with a boiling point of less than 204° C. In the second row, the temperature of 343° C. as the End Point is input and the number of the cuts is again 1, showing that there is one cut between 204° C. and 343° C. The same procedure is used to input VGO and the residue. When the blend is submitted, a number of hypothetical components are automatically created, each representing one oil fraction.

Sulfur and Nitrogen Compounds

The other hypothetical components that are defined separately are sulfur and nitrogen compounds. These two are the base components for HDS and HDN reactions respectively. The sulfur and nitrogen compounds, found in each cut, are intrinsically different. However, these molecules are classified into two different classes; the first one including those present in the gas oil cut with boiling points of 300-600° C. and the second one including those in the residue fraction. Therefore, one sulfur compound and one nitrogen compound is defined in each of these two classes and HDS and HDN reactions are based on these hypothetical components. The average properties of such molecules, i.e. density, molecular weight, etc., are used to define such compounds.

Other Compounds

Other compounds that were added to the component list are hydrogen, water, hydrogen sulfide (H₂S) and ammonia (NH₃) which either are reactants or are present in the system.

Components of oil are input to the simulator. Table 2 shows these components. The names are the simulator's default and can be change arbitrarily:

TABLE 2
The list of components defined in the simulation
Name	Definition	Boiling Point Range
NBP184	Naphtha	Below ° C.
NBP296	Middle Distillates	204-343° C.
NBP441	VGO	343-538° C.
NBP829	Residue	Over 538° C.
S-hydro	Sulfur compounds in non-residue oil	Below 538° C.
N-hydro	Nitrogen compounds in non-residue oil	Below 538° C.
S-residue	Sulfur compound in residue	Over 538° C.
N-residue	Nitrogen compound in residue	Over 538° C.

Horizontal Well Simulation

As mentioned, the mixture of oil/steam enters the horizontal wellbore via the perforations on its surface. These perforations are placed on the horizontal casing at intervals of about 15 cm (l′). An approximation of the total number of such perforations is calculated as:

$n=\frac{L_T}{I_i}=\frac{1000m}{0.15m}=6667$

Assuming a production rate of 100 m3/day, the fluid flow rate

$\left(V_i=\frac{V_T}{n}\right)$

for each of these perforations is calculated as 0.015 m3/day.

To simulate the wellbore, the wellbore was divided into 5 segments (an arbitrary number) and 5 corresponding feed entrances which are demonstrated in FIG. 6. The new feed stream entering each segment represents the combination of the feed streams that drain into the successive pipe segment via its perforations. The length of each segment is 200 meters and assuming an interval of 15 cm for these perforations, their total number for each segment is

$\frac{6667}{5}\cong 1333.$

Therefore the total flow rate of such a number of perforations is:

$Q=0.015\times 1333=20.0\frac{m^3}{\mathrm{day}}=0.83\frac{m^3}{h}.$

This flow rate corresponds to the new feed stream before each segment.

As shown in FIG. 6, the first feed stream enters the first pipe segment. Having passed through the first pipe segment, the first feed stream mixes with the second feed stream that enters the wellbore at the start of the second segment. The mix stream is directed to the second segment. This procedure continues until the final stream is ready to enter the vertical wellbore.

The feed consists of oil and water with an oil/water ratio of 2. In the simulation model, two streams were defined for oil and water separately to provide control over the model feed. The 5 feed streams in FIG. 6 are mixtures of oil and water streams that have a specific oil/water ratio of 2. Calculations showed that for a total production rate of 100 m3/day (standard ideal volume flow), each oil stream entering has a standard flow rate of 0.28 m3/h which sums to 1.4 m³/h (33.6 m³/day) for the 5 streams. Each water stream has a standard flow rate of twice as much as the oil or 0.56 m³/h.

Simulation Objects

The horizontal production wellbore was simulated in HYSYS using a pipe segment and mixers as shown in FIG. 7. The mixers do not influence the parameters of the system such as the pressure drop. For each pipe segment there is an energy stream by default which is automatically calculated based on the heat transfer input data and correlations. The pipe sizing is based on the information input by the user. This information determines the inner diameter and pipe material and then based on pipe schedules in the HYSYS database, the other data are determined. Table 3 shows some of the pipe segment input data:

TABLE 3
Pipe Segment Information
	Pipe flow correlation	Beggs and Brill
	Pipe material	Cast iron
	Ambient medium	Ground
	Ground type	Wet sand
	Buried depth	150	m
	Inner diameter	146.3	mm
	Outer diameter	168.3	mm
	Down-hole pressure	3	MPa
	Ambient temperature	200°	C.

Typical heat transfer data input by the user was as follows:

TABLE 4
Heat Transfer Information For Pipe Segments
Insulation type     No insulation
Thickness           0.01 m
Ambient medium      Ground/Wet sand
Buried type         150 m
Ambient temperature 200° C.

Pressure Drop

Pressure drop in the horizontal section was calculated using Darcy-Weisbach friction factor:

$\Delta P=f\times \frac{\rho {\stackrel{\_}{V}}^2L}{2D}$

where V is the velocity of fluid and is calculated as:

$V=\frac{Q}{\pi \frac{D^2}{4}}\left(Q=100m3\text{/}\mathrm{day}\mathrm{and}D=0.15\mathrm{cm}.\right).$

Therefore: V=0.06 m/s

HYSYS calculated the kinematic viscosity of the stream as 0.37 cSt. Therefore the Re number is calculated to be Re=24324 from:

$\mathrm{Re}=\frac{\mathrm{VD}}{v}$

To obtain the friction factor, Prandtl's friction factor correlation for smooth pipes was used:

$\frac{1}{\sqrt{f}}=2.0\log \left(\mathrm{Re}\sqrt{f}\right)-0.8$

Substituting the Re number into Prandtl's equation, results in an f of 0.025.

Thus, the final pressure drop due to the friction on the walls of the wellbore is calculated for a horizontal well of 1000 meters to be 300 Pa which is negligible.

The effect of scaling the length of the wellbore on the pressure and temperature profiles was investigated and three different models were created to show the effect of the number of pipe segments and feed locations. Table 5 shows these models.

TABLE 5
Three Different Models For Horizontal Wellbore Simulation
	Number of	Segment	Flow rate of each
	segments	length (m)	feed stream (m3/h)
	Model 1	5	200	0.83
	Model 2	10	100	0.42
	Model 3	20	50	0.21

The results show that the pressure drop does not change substantially in either of these models. Also the final temperature of the feed does not show a significant change (about 5° C. decrease for 1000 meters wellbore). Therefore the scaling used in this model (5 segments) was maintained as the base model.

Vertical Well Simulation

Upgrading reactions take place in the vertical section of the wellbore. The heating system, located at the start of the vertical well increases the temperature of the oil and water mixture. This increase is assumed to take place within the first 25 meters of the well. This length is arbitrary; however the exact length will be dictated by the power/intensity of the down-hole heat equipment. The heated section of the well is divided into 5 segments which are 5 meters in length. For hydrotreating reactions, each segment provides a temperature increase of 25° C. to eventually increase the stream temperature to 350° C. along the first 25 meters of the vertical wellbore. Similarly, for the hydrocracking reactions the temperature will be progressively increased depending on the final desired temperature. For a reacting temperature of 400° C., each segment must provide some 35° C. of temperature increase.

The vertical well is simulated as a number of plug flow reactors which are arranged in series. The reason for using more than one plug flow reactor is to allow for a higher level of control over the model parameters and to provide better tracking of the gradual increase in the conversions due to temperature and pressure changes along the wellbore. A schematic of the reactor simulation is shown in FIG. 8.

In this figure x_i is the conversion taking place in each reactor segment. The total conversion is given by:

x_T=1−(1−x₁)(1−x₂)(1−x₃)(1−x₄)(1−x₅)

where x_T is the total conversion.

FIG. 9 is the HYSYS process flow schematic for the vertical wellbore reactor. Oil, water and hydrogen streams may have different ratios and temperatures before entering the vertical wellbore. Stream 1 represents the feed entering the vertical well with a temperature that may be varied by the user (220° C. in this model). Section 2 shows five plug flow reactors, each 5 meters long, and their energy streams which control the outlet temperatures. Section 3 is the long plug flow reactor without an energy stream whose length may vary between 75 m and 475 m. Section 4 is the product stream.

Simulation Objects

The streams and the mixers in the vertical section are defined similar to those in the horizontal section. The major simulation object that the vertical section contains is the group of plug flow reactors.

The plug flow reactors in this section serve both as the vertical well and as reactors for hydroprocessing reactions. Each plug flow reactor is defined by 3 major sets of input data: the geometry, the reaction sets and the specific parameters such as pressure drop. The plug flow reactor can also have an energy stream when an understanding of the heat transfer parameters or temperature changes exist. For the geometry, the desired length and diameter of the reactor is input as is the value for the void fraction of the reactor. Each reactor can have one single or multiple sets of input reactions. For instance in the case of hydrocracking, four reactions in a network take place simultaneously. However in hydrotreating, only one reaction is active at a time. The pressure drop due to the hydrostatic head must be independently calculated and input.

Pressure Drop

The pressure drop in each segment is mainly due to hydrostatic head and not friction. The pressure drop due to the friction is calculated for the vertical wellbore (length 200 m) from:

$\Delta P=f\times \frac{\rho {\stackrel{\_}{V}}^2L}{2D}$

where the velocity is calculated as:

$V=\frac{Q}{\pi \frac{D^2}{4}}\left(Q=100m3\text{/}\mathrm{day}\mathrm{and}D=0.15\mathrm{cm}.\right).$

Therefore: V=0.06 m/s

Similar to the calculations for the horizontal section, assuming a kinematic viscosity of 2 cSt, Re number is calculated as 4900 and the friction factor is: f=0.038 from Prandtl's equation.

Based on the calculated values, the pressure drop due to the friction in the wellbore is negligible (171 Pa). Therefore, the total pressure drop will be due to the hydrostatic head pressure drop only:

ΔP=ρgh

For one segment of 5 meters, the static loss is 37.5 kPa, when ρ=750 kg/m3.

Reactions

The hydrotreating and hydrocracking reaction paths and the kinetic values are based on existing literature data. The power law form of the rate equation was chosen as the basis for this study.

r=kC₁^mC₂ⁿ

where k=k₀ exp(−Ea/RT), C₁ is the concentration of sulfur or nitrogen compound and C₂ is the concentration of hydrogen.

The data that must be found are the order by which the reactants take part in the reactions (m and n) and the values of k₀ and Ea for each reaction. To obtain the most proper reaction paths and kinetic data, two sources where chosen for simulation kinetic model, namely hydrotreating and hydrocracking reaction data.

Hydrotreating Reactions

In hydrotreating, there are two reactants and two products. In HDS, the sulfur compound and the hydrogen are the reactants and a desulfurized hydrocarbon and H₂S are the products. The following is the general form of an HDS reaction, where S_Comp is the sulphur compound and a and b are the stoichiometry coefficients:

S_Comp+aH₂→HC+bH₂S

In the case of HDN the nitrogen compound and the hydrogen are the reactants and a denitrogenized hydrocarbon and NH₃ are the products:

NComp+a′H2→HC+b′NH3

The order of the reactants in the reactions (m and n) and the values of k₀ and Ea for each reaction was based on literature reported values for HDS and HDN reactions as shown in Table 5.

TABLE 5
Kinetic Data For HDS And HDN Reported By Ferdous Et Al. And
Used In This Model - Low Boiling Point (<538° C.) Fractions
	Order	k0 (1/h)	Ea (kJ/mol)
	Scomp <538° C.	1.5	2.7E7	87
	Ncomp <538° C.	1	1E6	74

For residue compounds, the kinetic constants are different. For this study, no proper set of data for HDS and HDN reactions of residue oil, with boiling points of over 538° C. was found. Therefore, an approach reported by Trytten et al. was used to calculate the frequency factor and activation energies of residue sulfur and nitrogen compounds based on the corresponding values in the fractions below 538° C. In this approach the authors showed that the rate constants for HDS and HDN reactions decreases logarithmically with increasing the feed average molecular weight. The following equation was derived from their figures to estimate the rate constant for the residue compounds, where an approximation of their molecular weight is available:

log kR−log k1=−5.3(log MWR−log MW1)

where k1 is the known rate constant of a specific fraction, MW1 is the corresponding molecular weight, kR is the unknown rate constants of residue compounds and MWR is an approximation of the residue molecular weight.

Based on this correlation, an approximate value for both HDS and HDN rate constants of the residue compounds is obtained which is summarized in Table 6. The activation energy is assumed to be the same as the lighter fraction ones:

TABLE 6
Frequency Factor For HDS And HDN Used In This Model -
High Boiling Point (>538° C.) Fractions
k0 (1/h)
Scomp >538° C. 1.6E5
Ncomp >538° C. 1E4

Hydrocracking Reactions

The reactants in the hydrocracking reactions are a hydrocarbon and hydrogen. The product is a lighter hydrocarbon of low molecular weight. The following formula represents the general form of a hydrocracking reaction where HCl represents the heavy hydrocarbon, HC2 is the lighter one and a is the stoichiometry coefficient for hydrogen:

HCl+a H₂→HC2

For instance, HCl can be the residue and HC2 can be the vacuum gas oil (VGO).

For hydrocracking, the order of the hydrocarbons participating in the reaction was considered to be 1. The hydrogen order was assumed to be zero, assuming that the hydrogen is present in excess and at a high partial pressure.

As previously mentioned, each hydrocarbon family cracks into lighter hydrocarbons through reacting with hydrogen. In other words, the products of hydrocracking of heavier molecules are the reactants of another set of hydrocracking reactions which cracks these into lighter components. As a result, these reactions are not independent and occur through a network of reactions. Sanchez et al. suggested the network as shown in FIG. 10.

Calculated values of k₀ and Ea for the reactions paths of FIG. 10 showed that some of the reactions in this network had very small frequency factors and practically did not proceed to considerable conversions. Eliminating the low conversion reactions, a simplified form as shown in FIG. 11 was derived:

As can be seen from FIG. 11, there are four major reactions that take place in hydrocracking of oil which are hydrocracking of residue to VGO, hydrocracking of VGO to middle distillates, hydrocracking of residue to middle distillates and hydrocracking of residue to naphtha. The kinetic data for conventional catalysts and for all four of these reactions are presented in Table 8.

TABLE 8
Kinetic Data For The Simplified Hydrocracking Network
	k0 (1/h)	Ea (kJ/mol)
Path 1	7.4E14	202.7
Path 2	3.3E11	165.1
Path 3	4.8E12	184.8
Path 4	3.7E10	158.8

It is noted that the residue kinetic data for HDS, HDN and HyCr are different when they are processed as pure components rather than diluted in lighter fractions. When diluted, the viscosity and diffusion constraints within the oil are reduced resulting in higher conversions.

Hydrogen Pressure Effect

The hydrogen partial pressure effect was considered in the final conversion extent. The kinetic data for the conventional catalyst were obtained from the literature where the hydrogen pressure is higher than the one used in this study (3 MPa). For both hydrocracking and hydrotreating reactions, usually but not always, at higher partial pressures of hydrogen higher kinetic constants and consequently higher conversion can be expected and may require correction. In those cases where such a correction is not trivial, hydrogen consumption in the reactions at various hydrogen pressures can be an indication of the change in the conversion level which provides a similar prediction for the conversion drop.

It should also be noted that if the steam is assumed to be present in the reaction medium, the hydrogen partial pressure will be even lower than 3 MPa depending on the amount of steam present. In severe cases it can lower the hydrogen partial pressure to an extent that the conversion drops to zero.

The kinetic data for the ultra dispersed catalysts (discussed below) did not require the corrections applied to the conventional catalyst data of the literature; the reason being that the UD catalyst data are obtained at pressures near the simulation model conditions.

Reaction Simulations

The reactions may be defined in various forms such as inputting the kinetic data in power law or Langmuir-Hinshelwood forms, introducing the conversion function form based on the temperature or introducing the function form of the equilibrium constant based on free Gibbs energy.

Three sets of information are completed to develop each reaction model. The first set contains the stoichiometry information as well as the order by which each component participates in the reaction rate equation. Note, the reaction order for HDS reactions is 1.5 and for HDN reactions is 1. Hydrogen consumption in hydroprocessing reactions was used as an indication of the relative volume or mole numbers of reactants (oil fractions and hydrogen). Using the hydrogen consumption information for the reactions of interest, the corresponding stoichiometry number by which hydrogen takes part in each reaction was obtained. A summarized example of HDS reactions data is presented in Table 9:

TABLE 9
HDS Reaction Stoichiometry and Orders
Component	Stoich Coeff	Forward Order	Reverse Order
H2	−7	0	0
HC	−1	1.5	0
Desulfurized HC	1	0	0
	(Balanced)
H2S	0.414	0	0

where HC is the hydrocarbon with sulphur in its structure.

Table 10 shows an example of another set of information required to define a reaction in HYSYS. These information are:

• • The ‘Base Component’ which is the limiting reactant in the reaction medium;
  • The units for the reactants that take part in the reaction rate equation; and,
  • The ‘Basis’ which is the form of the ‘Basis Component’ as input into the model which can be mass fraction, mass concentration, mole fraction or mole concentration.

The reaction phase which is one of the following cases: liquid, vapour, overall (mixture of both liquid and vapour). The reaction phase for hydroprocessing reactions is normally liquid phase, where the products will partly join the vapour phase and exit the reactor rather quickly.

TABLE 10
Additional Information For Hydroprocessing Reactions
Defined In The HYSYS Model
Basis          Mass Fraction
Base Component HC
Reaction Phase Liquid
Basis Units    No units (wt fraction)
Rate Units     kgmole/m3-h

Knowing that any reaction is the result of both forward and reverse reactions happening at the same time, the general form of a kinetic reaction is:

r=k×f(Basis)−k′f′(Basis)

where k is the kinetic constant for the forward reaction and is defined as:

k=A×exp{−E/RT}×T^β

and k′ is the kinetic constant for the reverse reaction:

k′=A′×exp{−E′/RT}×T^β′

T is temperature in Kelvin. E and E′ are the activation energies for the forward and reverse reactions respectively.

The final tab in the reaction section inputs the frequency factor, the activation energy and the β factor. This factor shows the dependence of k₀ on temperature and is zero in most cases. An example of such a table for HDS reaction of non-residue fraction molecules is shown in Table 11:

TABLE 11
Inputting Frequency Factor, Activation Energy And Temperature
Dependency Factor For Kinetic Constants
A 5.4E8 h-1
E 85 kJ/mol
β 0

Results & Discussions

The results include those for hydrotreating including both HDS and HDN reactions, hydrocracking using conventional catalyst in the presence and absence of water and finally hydrocracking of heavy oil using ultra dispersed (UD) catalyst. Conventional catalyst kinetics was found in the literature for the commonly used catalysts as discussed above. UD catalyst kinetics are discussed below.

The hydrotreating results show the conversion percents of sulfur and nitrogen compounds and final changes in their weight percent due to the reactions at different temperatures and residence times.

In the case of hydrocracking, the volume percent change at various residence times and temperatures for all the fractions are shown together with the increase in the API gravity of the oil, which is a primary goal of upgrading. The kinetics of reactions using UD catalysts are shown. In addition, the API gravity increase in the oil using two different catalysts is compared.

Hydrotreating using Conventional Catalyst

The results of hydrotreating are presented at various temperatures and residence times. Where steam/oil ratio (SOR) is zero and the wellbore diameter is kept constant (0.15 m), the residence time only depends on the length of the wellbore which changes in the simulation runs. The production rate of oil is constant (1.39 m³/h), and it is assumed that the water is separated from the system. The residence time is simply calculated by dividing the volume of the reactor by the production rate.

The results are presented versus the wellbore length rather than the residence times to provide a better understanding of the physical requirements for these reactions. Table 12 can be used as a reference to compare the residence times and Liquid Hourly Space Velocities (LHSV) corresponding to each reactor length for the aforementioned conditions:

TABLE 12
Residence Times And LHSV Corresponding To Each Wellbore
Length (Diam. 15 Cm - Production Rate 1.39 m3/h)
Wellbore Length (m)	Residence Time (h)	LHSV (h-1)
100	1.27	0.79
200	2.54	0.39
300	3.81	0.26
500	6.35	0.16

For the residence times shown in Table 12, the conversion extents for the HDS reactions of non-residue molecules were obtained through the simulated plug flow reactors. FIG. 12 shows the results of HDS reactions of the non-residue lumped fractions (boiling points below 538° C.) and indicates the HDS conversion increases with wellbore length. Also comparing the three simulations, higher temperatures result in higher conversions.

The conversion extent for HDS reaction on the residue fraction was also modeled. As previously mentioned, the frequency factor reported in the literature for residue molecules is significantly lower that that of non-residue molecules. Therefore, it was expected that the conversion extent would be lower for such molecules compared to non-residue ones. FIG. 13 shows the conversion percents of HDS of the residue fraction and confirms the expected trend. FIG. 13 also indicates that HDS of residue compounds does not occur to a considerable extent. For a 500 meter wellbore, the conversion at 375° C. is about 4% where the similar value for non-residue lumped fractions is as high as 95%.

Similar simulations were performed for HDN reaction of heavy oil. Again the results are plotted against the wellbore length at different temperatures. The conditions are the same as those in HDS reactions (diameter is 0.15 m and productions rate for oil is 1.39 m3/h). FIG. 14 shows the conversion extent for non-residue nitrogen molecules in the heavy oil. The trend is similar to that of HDS conversion and the conversions are higher for higher temperatures and residence times.

The simulation model for the residue fractions in HDN reactions was also modeled (FIG. 15). It was noted that considering the low partial pressure of hydrogen, the HDN reactions have low conversion values at the experiment temperatures. As previously mentioned, the frequency factor for residue nitrogen molecules is as low as 1×104 h⁻¹. The small magnitude of this value which is two orders of magnitude smaller than that of non-residue compounds (1×106 h⁻¹), justifies the inconsiderable conversion of the residue compounds at similar conditions. FIG. 15 indicates that the values are lower than those of the HDS. The maximum point is for HDN at 375° C. and 500 m well which is only 0.4% conversion.

FIG. 16 compares the results of HDS and HDN reactions at 350° C. and different residence times for non-residue lumped fractions which shows the considerably higher conversion percent for the HDS reactions of the non-residue lumped fractions in comparison with HDN reactions. The results of HDS and HDN reactions on residue fraction are also compared in FIG. 17.

When the conversion percent was determined by the simulation model, the sulfur percent of feed and product streams can be compared. A simple back calculation from the conversion formula provides the product sulfur percent:

S_p=S_f−X×S_f

where S_p is the sulfur percent of the product stream and S_f is the sulfur percent of the feed stream and X is the percent conversion. Based on this calculation, the percent sulfur in the feed and product streams can be compared. FIG. 18 compares the percent sulfur at different wellbore lengths at 350° C.

It can be seen that longer wellbores provide deeper hydrodesulfurization. Also as the figure suggests, the decrease in the product's sulfur percent is mainly due to the HDS of non-residue compounds rather than residue compounds. FIG. 18 also shows that the maximum global HDS for 500 meters of wellbore at a reactor temperature of 350° C. is some 30.6%.

Similar calculations provide the nitrogen percent change in the oil sections. FIG. 19 shows the percent conversion for hydrodenitrogenation for different oil cuts.

Similar to FIG. 18, most of the percent nitrogen decrease in the products of HDN is due to the HDN of non-residue lumped fractions rather than residue fraction. Nitrogen weight percent almost does not change in the residue for any length of wellbore. The global HDN for 500 meters of wellbore at 350° C. is 13.6%.

Hydrocracking using Conventional Catalyst

The major goal of conventional hydrocracking is obtaining light fuels such as diesel and gasoline. Hydrocracking of bitumen targets the production of a lighter crude oil through down-hole upgrading. Thus, the produced oil should have a lower molecular weight than that of bitumen. Through hydrocracking, oil fractions react with hydrogen in the presence of a catalyst and crack into lighter molecules. This continuous trend of conversion, gradually changes the proportions of the original oil fractions by reducing the volume percent of the heaviest fractions and increasing the lighter fractions.

A study of the volume percent change in various fractions of the oil, as a result of hydrocracking was undertaken. The independent variables were temperature, residence time and steam/oil ratio (SOR). It should be noted that the results may have a percent error when SOR is not zero, as the presence of water in the system may lower the frequency factor in the kinetics of hydrocracking reactions and consequently lower the extent of conversions. An extensive search of the available literature data was conducted to obtain information regarding this effect on the kinetics of reactions; however no proper data was found. Therefore, in this study there are only two effects that the presence of water imposes to the system: increasing the velocity of fluids in the wellbore (decreasing the residence times) and decreasing the concentration of reactants. Both of these effects are detrimental to the conversion extent.

As mentioned above, hydrocracking consists of a network of reactions that occur simultaneously. In this network, residue hydrocracks into VGO, middle distillates and naphtha. Also VGO hydrocracks into middle distillates. Therefore the volume percent of residue is always declining while the middle distillates increase. In the case of VGO, the increase or decrease of volume percent depends on the extent to which residue conversion proceeds and if this will be higher than that of the VGO conversion into middle distillates. In most cases, as the results show, the volume percent of VGO will increase along with an increase in the quantity of middle distillates. The reason is that kinetic parameters of residue hydrocracking into VGO promote a faster reaction than conversion of VGO into middle distillates. In other words, there is an accumulation of VGO in the process of production and conversion, resulting from the faster production rate in comparison with the conversion rate. For hydrocracking and for further comparisons, the composition of feed oil is shown in FIG. 20.

Steam/Oil Ratio: 0

For hydrocracking using conventional catalyst in the absence of water, the hydrogen flow rate is 15 kgmol/h and the oil flow rate is 1.389 m³/h. The residence times corresponding to each wellbore length, is shown in Table 12.

FIGS. 21 to 24 present simulation model results for hydrocracking at four different temperatures, namely 425° C., 350° C., 375° C. and 403° C. respectively.

FIG. 21 shows the volume percent increase of middle distillates with an increase in the wellbore length (residence time). As shown, the VGO increases by up to 68% at a length of 300 meters and then suddenly decreases to 60%. The reason is that the amount of residue at 300 meters is so low that its rate of conversion no longer exceeds that of VGO at a point between 300 and 500 meters. This result implies that VGO from that point on will be converted into middle distillates and there is no net production of VGO anymore.

Another result observed in this figure is the constant amount of naphtha fractions. It can be seen that the volume percent of this fraction is not increasing when the depth of wellbore increases to 500 m. This result suggests that the conversion of residue, as the only reactant for production of naphtha is very limited and this reaction can not compete with the other two reactions that residue undergoes (hydrocracking into VGO and middle distillates). This figure is produced at the highest temperature that the simulations in this work are carried out (425° C.). At lower temperatures, the conversions of these irreversible reactions are expected to be lower and therefore naphtha conversion will not increase which was observed through the simulation experiments; therefore, in FIGS. 22-24 naphtha conversion is not shown:

FIG. 22 shows that hydrocracking conversions at 350° C. with a conventional catalyst are very small that there are no major changes in the volume percents at different wellbore lengths. This is due to the low temperature level which is not able to promote the reactions to considerable extents.

FIG. 23 shows the conversions at 375° C. The volume percents change from 100 m to 500 m. The trend of this change, as was mentioned earlier, is that the volume percent of middle distillates and VGO increase when residue decreases. FIG. 24 shows the same trend as FIGS. 22 and 23, however the conversions are higher due to the higher temperature.

Comparing these four figures shows that the middle distillates increase in volume percent through hydrocracking reactions, VGOs may increase or decrease depending on the amount of residue present and finally residues always decreases. It was also noted that an increase in residence time increases the conversions in each of the figures by increasing the wellbore length. The effect of temperature increase can also be observed by comparing FIGS. 21-24; however for a clearer view the conversions are plotted vs. temperature at specific lengths in FIGS. 25-28. Note the wellbore diameter is 15 cm and production rate is 1.39 m³/h and that conversion increases gradually as the simulation model temperature increases.

Similar figures were produced for wellbore diameter of 0.1 m. This allows for a better comparison of the results over a broad range of residence times. The residence times corresponding to each wellbore length are given in Table 13.

TABLE 13
Residence times and LHSVs corresponding to each wellbore
length (diameter 10 cm - production rate 1.39 m3/h)
Wellbore Length
(m)	Residence Time (h)	LHSV (h-1)
100	0.56	1.77
200	1.13	0.88
300	1.70	0.59
500	2.83	0.35

Conditions such as SOR are the same as the previous set. FIG. 29 shows the volume percent change at 425° C. for all three fractions of oil and the residue. Similar to FIG. 21, the naphtha volume percent does not change with length. FIGS. 30-32 show the results of runs at 350° C., 375° C. and 403° C. respectively. FIGS. 33-36 show the results of runs for different length well bores. The trend of change in the fractions is the same as the previous set presented for diameter of 15 cm; however the conversions are lower because of small diameter or small residence time. The volume percent of residue decreases while that of the VGO and middle distillates increases. For 350° C. the conversion is so low that no considerable change in the volume percents is observed.

Steam/Oil Ratio: 1

Results for hydrocracking at a SOR 1 show that the flow rate of water is 1.39 m³/h; equal to the oil's. Hydrogen flow rate is as the previous case and the simulation runs are again based on the conventional catalyst. The residence times based on the wellbore length are given in Table 14:

TABLE 14
Residence times and LHSVs corresponding to each wellbore
length (diameter 15 cm - production rate 2.78 m3/h)
Wellbore Length (m)	Residence Time (h)	LHSV (h-1)
100	0.64	1.57
200	1.27	0.79
300	1.91	0.52
500	3.18	0.31

The results for the volume percent change are presented vs. the wellbore length in FIGS. 37-40 and later vs. temperature in FIGS. 41-44:

The trend in FIGS. 37-44 is similar and shows that the residue decreases and VGO and middle distillates increase. For a temperature of 350° C., the conversions are almost zero and for higher temperatures they increase. The results of a SOR 1 are also shown vs. the temperature to show the effect of wellbore length at a constant diameter of 15 cm. This data also shows a similar trend to the previous cases meaning that higher temperatures can significantly increase the conversions.

Steam/Oil Ratio: 10

Hydrocracking at a SOR 10 was also evaluated as part of this study. The residence times are shown in Table 15 and results shown in FIGS. 45-48.

TABLE 15
Residence times and LHSVs corresponding to each wellbore
length (diameter 15 cm - production rate 15.27 m3/h)
Wellbore Length (m)	Residence Time (h)	LHSV (h-1)
100	0.12	8.65
200	0.23	4.32
300	0.35	2.88
500	0.58	1.73

API Gravity Increase

The API gravity increases by a decrease in the specific gravity of oil as shown:

$\mathrm{API}=\frac{141.5}{{\mathrm{SG}}_{\left(\mathrm{at}60°F.\right)}}-131.5$

where SG is the specific gravity of the component at 60° F. Any increase in the density of oil, results in a specific gravity increase which then results in an API gravity decrease.

To calculate the specific gravity of the mixture the following formula is used:

$\mathrm{SG}=\sum _i{\mathrm{Vol}\mathrm{fr}}_i\times {\mathrm{SG}}_i$

Table 16 shows the values for specific gravities of the oil cuts, based on a typical Alberta bitumen assay:

TABLE 16
Specific Gravity Of Oil Fractions
Fraction           Specific Gravity
Naphtha            0.78
Middle Distillates 0.9
VGO                0.96
Residue            1.06

By assuming that the specific gravity values do vary little with conversions, the specific gravity of the product stream is calculated based on the volume percents of the oil cuts. Then, using the value of the specific gravity of the oil, the API gravity is obtained and can be compared to that of the feed.

FIGS. 49-51 present the API gravity increase as a result of the volume percent change in the fractions at three different SORs. These figures show that any increase in the residence time results in an increase in the API gravity of the feed. Note the higher temperatures provide higher API changes. When the SOR is 10, the steam/oil ratio is very high, such that the API gravity change is zero in all cases except for a temperature of 425° C.

As down-hole separation of steam from the oil is not a trivial process, economic evaluations should be conducted to compare the API gravity increase in the oil at various SORs and investigate if the differences are encouraging to separate the steam down-hole and if so to what extent. FIGS. 49-51 compare the API gravity at 425° C. and 403° C. for a SOR of zero, 1 and 10.

As the figures show, the API gravity increases for SOR of zero is slightly higher than that of SOR 1. The comparisons show that a SOR of 1 results in a lower API increase. An explanation would say that the dilution of the hydrocarbons in the presence of water reduces the API gravity change. Higher amounts of water result in a lower concentration of reactant and consequently lower conversions as shown in FIG. 51. FIGS. 52-53 show API gravity increase at 425 C and 403 C for SOR of zero and 1.

Hydrocracking Using Ultradispersed Catalyst

The kinetics of hydrocracking reactions using ultra dispersed (UD) catalyst are presented and compared with those of conventional catalysts. To obtain UD kinetics, the experimental results of hydrocracking of a sample of Peace River bitumen were used. The activation energy of the UD catalyst is assumed to be the same as that of the conventional catalyst; however, the frequency factors are higher for the UD catalyst due to higher available surface area. These k₀ values are shown below.

To obtain the kinetic values of the UD catalyst, a method to estimate these values based on the reaction conversions provided from laboratory experiments was employed. The experimental results of hydrocracking on Peace River bitumen provide the conversions of residue, VGO and middle distillates at a specific residence time. The simulation model was run at various k₀ values corresponding to each reaction path in the hydrocracking network to find a new k₀ which would produce a similar conversion to the experimental data. In other words, by trial and error new k₀ values were found that produced heavy oil with conversions similar to those that the UD catalyst produces through the simulation runs in accordance with the hydrocracking reaction network presented above.

Table 17 compares the frequency factors of each reaction:

TABLE 17
Comparing the Frequency Factors for Conventional
Catalyst and UD Catalyst
	k0 (1/h)	k0 (1/h)
Reaction Path	Conventional Catalyst	UD Catalyst
1	7.4E14	5.0E15
2	3.3E11	1.9E12
3	4.8E12	9.6E12

It should be noted that similar to what the literature data for conventional catalysts showed, the k₀ in the conversion of residue to naphtha for UD catalysts in reaction path 4 is negligible when compared to that of paths 1, 2 and 3.

By comparing the kinetics of UD catalyst with a conventional catalyst, in Table 17, it can be seen that the UD catalyst have higher k₀ values for hydrocracking reactions. This means that such reactions will have higher conversions using this kind of catalyst.

The simulation model was run based on the kinetics of the UD catalyst and was compared to those for the conventional catalyst. FIG. 54-57 shows the results at 425° C., 350° C., 375° C. and 403° C. respectively. FIG. 53 shows that the naphtha volume percent does not change. This is also the trend for all the other temperatures. In other words naphtha volume percent stays constant. FIG. 58 shows the API gravity increase due to hydrocracking using UD catalyst at various temperatures of 350° C., 375° C., 403° C. and 425° C.:

FIGS. 59-61 show the API gravity increase for the UD catalyst compared to that of conventional catalyst and as expected shows a higher conversion and therefore a higher API gravity increase.

Simulation Results Summary

A summary of the simulation results is shown in Table 18.

TABLE 18
Simulation Results Summary
	Simulation Conditions	Result
HDS (Hydrogen	residence time of 50 min @	over 2 wt % decrease in the
Desulpherization):	350° C.; water separated;	sulfur content of the crude
	200 m	heavy oil (2 wt % decrease in
		the distillate, 1.7 wt %
		decrease in the residue)
HDN (Hydrogen	residence time of 50 min @	over 0.03 wt % decrease in the
Denitrogenation)	350° C.; water separated;	nitrogen content of the crude
	200 m	heavy oil
Hydrocracking	Residence time of 50 min @	over 3.5 degree API change
	423° C., water separated;
	200 m
Hydrocracking	residence time of 50 min @	change in the volume percent
	423° C.; in the presence of	of the three main fractions if
	water	heavy oil is: Middle
		Distillates: 2% increase -
		Vacuum Gas Oil: 3% increase -
		Residue 5% decrease
HDS/HDN	residence time of 75 min @	heavy oil fractions with
	350° C.; water separated	boiling points of below 540° C.:
		80 mol % conversion in the
		sulfur content in the distillate
		fraction, 42 mol % conversion
		in the nitrogen content of the
		distillate fraction

Based on the results generated by the simulation, a partially upgraded crude oil product with lower amounts of contaminants and enhanced transporting properties was produced.

Methodologies and Catalyst Compositions

In a preferred embodiment, the catalysts applicable to the invention are continuously introduced in the form of a micro-nano particulate dispersed in hydrocarbon media as described in Applicant's co-pending application (U.S. application Ser. No. 11/604,131 and incorporated herein by reference) or a conventional catalyst or catalyst system such as a fixed bed catalyst is used.

Different catalyst compositions may be introduced simultaneously within different hydroprocessing zones. For example, catalysts formulated to enable hydrogenation, hydrotreating including desulfurization, hydrodemetalization and denitrogenation and hydrocracking reactions may be introduced into one or more hydroprocessing zones by means of separate or combined injection systems wherein the point of injection of the hydrogen and/or catalysts will define the location of one or catalytic hydroprocessing zones. The catalyst formulation and process conditions may be adjusted such that a suite of in-well processing elements can effect overall upgrading involving both hydrotreating and hydrocracking or other combinations of processes to produce an oil of desired quality and specification.

Temperature, hydrogen flow and catalyst flow may be continuously monitored and adjusted at the surface based on produced oil composition analysis at surface or using in situ sensors in order to adjust the downhole reactor variables during processing. Such adjustments may be made in real time using a variety of physical proxies for oil properties including viscosity, bulk chemical composition, SARA (Saturates, Aromatics, Resins, Asphaltenes) and/or chemical proxies indicative of the reaction regime (hydrotreating, hydrocracking, visbreaking etc).

In accordance with one embodiment of the invention, catalyst compositions as described in Applicant's co-pending United States application, are useful in in situ hydrocarbon upgrading applications. Catalyst compositions, characterized by their particle size and ability to form microemulsions are described herein.

The catalyst compositions are bi- or tri-metallic compositions dissolved in a protic medium containing a VIII B non-noble metal and at least one VI B metal (preferably one or two) in the presence of a sulfiding agent. The atomic ratio of the Group VI B metal to Group VIII B non-noble metal is from about 15:1 to about 1:15. Suitable catalyst compositions can be used in a variety of hydrocarbon catalytic processes to treat a broad range of feeds under wide-ranging reaction conditions such as temperatures from 200° C. to 480° C.

Bi-metallic catalysts of the general formula:

B_xM_yS_{[(1.1 to 4.6)y+(0.5 to 4)x]}

where B is a group VIIIB non-noble metal and M is a group VI B metal and 0.05≦y/x≦15 are effective.

In more specific embodiments, 0.2≦y/x≦6 and preferably y/x=3.

A second class of catalysts described as tri-metallic catalysts of the general formula:

B_xM1_yM2_zO_{(2 to 3)z}S_{[(0.3 to 2)y+(0.5 to 4)x]}

where B is a group VIIIB non-noble metal and M1 and M2 are group VI B metals and 0.05≦y/x≦15 and 1≦z/x≦14 are also effective.

Further examples of tri-metallic catalysts include those wherein the y/x ratio is in the range of 0.2<y/x<6. The range z/x is preferably determined by the desired use of the catalyst. For example, selectivity to lighter hydrocarbons (C1-C5) will preferably have a z/x of 10<z/x<14 and more preferably z/x=12. Alternatively, selectivity to intermediate hydrocarbons for mild hydrocracking (Low cracking functionality) will favor 1<z/x<5 and preferably z/x=3.

Formula Examples

As examples, if y/x=0.05, y=1 and x=20. Thus, at this y/x ratio,

B_xM_yS_{[(1.1 to 4.6)y+(0.5 to 4)x]}

would include catalyst compositions ranging from B₂₀MS_11.1 to B₂₀MS_84.6. If y/x=15, y=15 and x=1, at this y/x ratio,

B_xM_yS_{[(1.1 to 4.6)y+(0.5 to 4)x]}

and would include catalyst compositions ranging from BM₁₅S₁₇ to BM₁₅S₇₃.

CONCLUSIONS

The results show that a hydraulic model for a heavy oil wellbore was successfully linked with the kinetics of hydroprocessing reactions. The hydraulic model showed that the pressure drop in the horizontal section was negligible (300 Pa for 1000 meters of wellbore). The reason for the low pressure drop is the low velocity of the liquids in the wellbore which mitigates high friction losses. In the vertical section, the friction loss was low again due to the low velocity. The major pressure drop in the system was due to the hydrostatic head of fluids in the vertical wellbore. Also, based on the hydraulic model it was also observed that temperature variations in the horizontal well are negligible.

Three kinetic models were successfully developed for the hydroprocessing of heavy oil that included hydrotreating using conventional catalyst kinetics and hydrocracking using conventional catalyst as well as UD catalyst kinetics.

It was shown via the simulation model that the major portion of hydrodesulfurization of heavy oil occurs through hydrodesulfurization of low boiling point sulfur compounds (under 538° C.) rather than the residue ones (boiling point over 538° C.). The reason was the substantially lower frequency factor that the residue compounds posses. At the longest length of the wellbore (500 m), which corresponds to a residence time of 6.35 h and at the temperature of 350° C. of the experiments, it was shown that the hydrodesulfurization conversion of residue cuts did not exceed 2% where this number for lighter fractions was 87%. At these conditions the total sulfur percent change in the product was 1.3 wt %. Also the global HDS conversion for the feed at the mentioned conditions (6.35 h residence time and 350° C.) was 30.6%.

It was also shown that hydrodenitrogenation reactions, similar to hydrodesulfurization, had higher conversions for the lighter fractions when compared to the residue compounds. The conversion extent for a wellbore with 500 m length (residence time of 6.35 h) and at 350° C., for residue nitrogen compounds was 0.2% and for the lighter nitrogen compounds was 60%. The total HDN resulted in 0.03% drop in the nitrogen weight percent. The global HDN conversion was 13.6% for these conditions (6.35 h residence time and 350° C.).

The results of hydrocracking simulation model showed that the residue volume percent decreased from the feed to the products whereas the middle distillates increased. It was also shown that naphtha volume percent stayed relatively constant under various conditions used in the hydrocracking simulation. The VGO volume percent showed either an increase or decrease depending on the amount of residue present in the system to produce VGO. Using the conventional catalyst kinetics, the simulation results showed that the API gravity of the oil in comparison with the feed increased about 9.7 points at a 500 m wellbore (residence time of 6.35 h) and at 425° C. and SOR of 0. At a SOR 1 the API gravity increased less than for the SOR 0 (about 1.3 points lower at a residence time of 3 hours and at 425° C.).

The hydrocracking results were also compared for two different catalyst kinetics. The results showed that the hydrocracking conversions were higher for the UD catalyst in comparison with a conventional catalyst. At 425° C. and residence time of 6.35 (wellbore length 500 m), the API gravity increase when using the D catalyst was 11.6 which is 1.9 points higher than when conventional catalyst kinetics was used. At 403° C., the API gravity increase using the UD catalyst kinetics was 8.7 which was 3.3 points higher than the increase for the conventional catalyst kinetics.

In summary, a new method for in-situ upgrading is provided that remedies many practical problems by using the wellbore capacity for the upgrading reactions. This approach optionally eliminated the need for down-hole catalyst placement by using ultra dispersed catalysts in liquid phase which enters the wellbore volume and gets produced with the upgraded oil. In this method, the necessary heat of the reaction is concentrated in the vertical wellbore rather than being introduced to the large down-hole bitumen reserve in order to minimize the heat loss. Finally the reactions occur over the ultra dispersed catalysts which offer a high effective contact area and have higher kinetic frequency factor for hydrocracking reactions which then results in substantially higher conversions.

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Claims

1. A method of upgrading heavy oil in a production well within a hydroprocessing zone comprising the steps of:

a. introducing a controlled amount of heat to the hydroprocessing zone;

b. introducing a selected quantity of hydrogen to the hydroprocessing zone to promote a desired hydrocarbon upgrading reaction; and,

c. recovering upgraded hydrocarbons at the surface.

2. A method as in claim 1 further comprising the step of introducing a catalyst to the hydroprocessing zone.

3. A method as in claim 2 wherein the catalyst is a nano-particle catalyst.

4. A method as in claim 3 wherein the catalyst is circulated within the hydroprocessing zone.

5. A method as in claim 1 wherein the hydroprocessing zone is a vertical section of a wellbore.

6. A method as in claim 1 wherein the method includes separating heavy oil from water prior to introducing heavy oil into the hydroprocessing zone.

7. A method as in claim 1 wherein the heavy hydrocarbon is bitumen and the upgraded hydrocarbons are characterized by an API gravity increase.

8. A method as in claim 3 wherein the catalyst is a bi-metallic catalyst of the general formula: BxMyS[(1.1 to 4.6)y+(0.5 to 4)x] where B is a group VIIIB non-noble metal and M is a group VI B metal and 0.05≦y/x≦15.

9. A method as in claim 3 wherein 0.2≦y/x≦6.

10. A method as in claim 9 wherein y/x=3.

11. A method as in claim 3 wherein the catalyst is a tri-metallic catalyst of the general formula: BxM1yM2zO(2 to 3)zS[(0.3 to 2)y+(0.5 to 4)x] where B is a group VIIIB non-noble metal and M1 and M2 are group VI B metals and 0.05≦y/x≦15 and 1≦z/x≦14.

12. A method as in claim 11 wherein the y/x ratio is in the range of 0.2<y/x<6.

13. A method as in claim 11 wherein 10<z/x<14.

14. A method as in claim 11 wherein z/x=12.

15. A method as in claim 11 wherein 1<z/x<5 and the upgrading process is mild hydrocracking.

16. A method as in claim 11 wherein z/x=3 and the upgrading process is mild hydrocracking.

17. A method as in claim 1 wherein the controlled amount of heat is introduced using any one of or a combination of electrical, hot fluid, or an in-well combustion device.

18. A method as in claim 1 wherein the production well includes at least two hydroprocessing zones and a different hydroprocessing reaction is controlled in each hydroprocessing zone.

19. A method as in claim 1 wherein the upgrading process is part of a steam flooding process including any one of steam assisted gravity drainage (SAGD), vapor extraction (VAPEX), cyclic steam stimulation (CSS) and CAPRI.

20. A method as in claim 1 wherein the upgrading reaction is hydrodenitrogenation.

21. A method as in claim 1 wherein the upgrading reaction is hydrodesulfurization.

22. A system for upgrading heavy oil in a production well within a hydroprocessing zone comprising:

a downhole heater for introducing a controlled amount of heat to the hydroprocessing zone;

a hydrogen delivery system for introducing a selected quantity of hydrogen to the hydroprocessing zone to promote a desired hydrocarbon upgrading reaction; and,

a surface recovery system for recovering upgraded hydrocarbons at the surface.

23. A system as in claim 22 further comprising a downhole water separator for separating water from heavy hydrocarbon, the downhole water separator operatively located upstream of the hydroprocessing zone.