Note: Descriptions are shown in the official language in which they were submitted.
WO 2015/143039 PCT/US2015/021258
System and Method for Controlling and Optimizing the Hydrothermal
Upgrading of Heavy Crude Oil and Bitumen
BACKGROUND
15
Technical Field
This invention relates to hydrocarbon processing, and more particularly to
systems and
methods for efficiently upgrading heavy crude oil.
Background Information
Introduction to Heavy Crude Oil
The average weight or density of crude oils extracted from oil fields globally
has been
increasing very gradually over time, a trend expected to continue
indefinitely. However, the
existence of large reserves of heavy and extra-heavy crude oils in some
countries means that the
as-produced weight of crude oil can increase much more rapidly on a regional
basis. Of
particular importance are the tar oils in the Orinoco Belt in Venezuela and
oil sand bitumen in
Alberta, Canada, which in aggregate are currently estimated as being 2-3 times
the size of the oil
reserves in Saudi Arabia. The density of Saudi Arabian crude oils, expressed
as API gravity or
.. API, may typically fall in the range of about 27 - 34 API, in the center
of which falls the
current global average. By contrast, the deposits in Venezuela and Alberta are
generally
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characterized as being heavy crude oils (HCO) or extra-heavy crude oils (EHCO)
for which the
corresponding densities may be regarded generally as being below about 22.3
API and about 10
API, respectively. (The lower the API value, the higher the density.) For
deposits that are
heavier still, such as in the case of some natural bitumen deposits in
Alberta, the term ultra-heavy
crude oil (UHCO) is sometimes applied. In most cases, the densities of native,
unmodified heavy
crude oils produced in Venezuela and Alberta are below about 15 API, and even
below 10
API. (Though the classification scheme used herein to differentiate crude oils
in terms of API
will be recognized by those skilled in the art, other conventions and criteria
exist, which may
apply different terms and API ranges and/or include other criteria such as
viscosity and percent
sulfur. Therefore, definitions used herein should not be regarded as limiting
but only illustrative.)
From the viewpoint of crude oil production and transport, HCO, EHCO, and UHCO,
the
entire group of which shall hereinafter be referred to inclusively as heavy
crude oils (HCO)
without limitation as regards exact composition or geological or geographic
origin, are
problematic because the same physico-chemical characteristics that cause their
elevated density
produce a corresponding increase in viscosity. By way of illustration that is
neither bound by
theory nor intended to be complete or applicable to all crude oils,
asphaltenes are a class of
diverse compounds known to affect density and viscosity directly and to have
concentrations in
HCO that are generally higher than in medium and light crudes. Having
molecular weights that
are high relative to other compounds in crude oils generally, increasing
asphaltene concentration
is generally accompanied by an increase in both density and viscosity. This
may be due to the
tendency of asphaltenes to self-associate, or it may be due to the formation
of dense microscopic
particles comprising a dense core of aggregated asphaltenes surrounded by
layers of other crude
oil components. Regardless of the mechanisms by which composition and
microscopic structure
cause elevated density and viscosity, HCO is generally not amenable to the
methods of
transportation and storage commonly applied to medium crude oils (about 22.3
API to about
31.1 API) and light crude oils (greater than about 31.1 API). For example,
if crude oil were
required to have a minimum API value of about 20 to be pipelineable, and if
transport by rail
tank car is precluded on the grounds of practical economics and logistics,
then delivery to market
of crude oil extracted from Albertan oil sands requires that it be somehow
upgraded to meet
pipeline specifications for density and viscosity.
Approaches to Upgrading Heavy Crude Oils
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Commercially relevant upgrading strategies currently applied in Alberta fall
into two
general categories. In the first, coking, hydrocracking, or other techniques
are applied to HCO to
chemically convert asphaltenes and other heavy components into lighter
materials, which are
recovered through distillation and blended to produce pipeline quality
synthetic crude oil. The
various conversion and recovery processes are related to those employed in oil
refining and the
overall approach is correspondingly capital intensive, adding an estimated $14
per barrel.
Furthermore, economic considerations preclude an implementation strategy
whereby
smaller-scale upgrading facilities may be located in or near numerous
production fields.
Producers therefore rely on another, simpler strategy whereby the bitumen and
heavy oil
are mixed with higher-value, lighter petroleum products at the wellhead to
produce diluted
bitumen (dilbit) that can be easily transported through pipelines. However,
several significant
issues are associated with dilbit. First, the diluent must be transported by
rail or pipeline to
production fields from distant refineries or gas processing plants where it is
produced. Second,
dilbit in pipelines may typically contain about 20% to 40% diluent,
effectively reducing the net
capacity of pipelines to carry unrefined crude. Compounding these issues, the
net cost for diluent
in terms of both the material itself and the facilities required to handle it
adds $10 ¨ $16 per
barrel of dilbit. However, beyond infrastructure and cost considerations looms
a broader
problem, namely, that diluent-based upgrading may not be a practical way to
meet future growth
of Canadian HCO production. Absent an alternative approach, Canada will be
required to import
ever increasing quantities of diluents. Currently, efforts are underway to
expand the pipeline
infrastructure from the Gulf Coast of the United Stated all the way to Alberta
via Illinois to carry
the "pentane plus" condensate by-product of shale gas production.
The need exists in the art for a new approach that requires lower initial
capital
investment, has lower ongoing operating costs, and combines the best features
of the two main
upgrading methods used currently: reduction of the density and viscosity of
the native crude
through conversion of asphaltenes and other heavy components into lighter
ones; and scalability
that permits distributed implementation at or near the wellhead to minimize or
eliminate the
reliance on diluent from remote sources.
SUMMARY
According to one aspect of the present invention, a system is provided for
upgrading a
continuously flowing process stream including heavy crude oil (HCO). The
system includes a
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fluid flow path configured to convey the process stream continuously
therethrough in a
downstream direction. The flow path includes a reactor configured to receive
the process
stream in combination with water, at an inlet temperature within a range of
about 60 C to
about 200 C. The reactor includes one or more process flow tubes defining an
aggregated
interior cross-sectional dimension transverse to the downstream direction, the
flow tubes
having a combined length of at least about 30 times the aggregated interior
cross-sectional
dimension. The reactor applies heat to progressively heat the process stream
to an outlet
temperature T(max)1 within a range of between about 260 C to about 400 C.
The reactor is
also configured to maintain the process stream at a pressure sufficient to
ensure that it remains
a single phase at T(max)1. A controller is configured to selectively adjust
the rate of flow of
the process stream through the reactor to maintain a total residence time in
the reactor of
greater than about 1 minute and less than about 25 minutes. This rate of flow,
in combination
with the flow tube length and cross-sectional area, and the progressive
application of heat, are
configured to minimize or prevent coke formation.
In another aspect of the invention, a method for upgrading a continuously
flowing
process stream including heavy crude oil (HCO) includes supplying the process
stream to the
aforementioned system.
The features and advantages described herein are not all-inclusive and, in
particular,
many additional features and advantages will be apparent to one of ordinary
skill in the art in
.. view of the drawings, specification, and claims. Moreover, it should be
noted that the
language used in the specification has been principally selected for
readability and
instructional purposes, and not to limit the scope of the inventive subject
matter.
BRIEF DESCRIPTION OF THE DRAWINGS
The present invention is illustrated by way of example and not limitation in
the figures of
the accompanying drawings, in which like references indicate similar elements
and in which:
Fig. 1 is a schematic diagram of a representative embodiment of a hydrothermal
heavy
crude oil (HCO) upgrading system of the present invention;
Fig. 2 is a schematic diagram of the embodiment of Fig. 1, with various
optional features;
Fig. 3 is a schematic diagram of the embodiment of Fig. 2, with further
optional features;
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Fig. 4 is a schematic diagram of the embodiment of Fig. 3, with additional
optional
features;
Fig. 5 is a schematic diagram of the embodiment of Fig. 4, with an additional
optional
feature;
Fig. 6A is a schematic diagram of the embodiment of Fig. 3, with an additional
optional
feature;
Fig. 6B is a view similar to that of Fig. 6A, with other optional features;
Fig. 7 is a schematic cross-sectional view of a reactor suitable for use in
one or more of
the embodiments of Figs. 1-6B, with temperature represented graphically
thereon;
Fig. 8 is a view similar to that of Fig. 7, of an alternate reactor;
Fig. 9 is a schematic diagram of another alternate reactor portion usable with
embodiments of the present invention;
Fig. 10 is a schematic diagram of still other reactor portions usable with
embodiments of
the present invention;
Fig. 11 is a schematic diagram of still other reactor portions usable with
embodiments of
the present invention;
Fig. 12 is a schematic diagram of an alternate embodiment of a hydrothermal
heavy crude
oil (HCO) upgrading system of the present invention;
Fig. 13 is a schematic diagram of yet another alternate embodiment of a
hydrothermal
heavy crude oil (HCO) upgrading system of the present invention;
Fig. 14 is a schematic diagram of another embodiment of a hydrothemial heavy
crude oil
(HCO) upgrading system of the present invention;
Fig. 15 is a schematic diagram of another embodiment of a hydrothermal heavy
crude oil
(HCO) upgrading system of the present invention;
Fig. 16 is a graphical representation of aspects of the embodiments of Figs.
12-15;
Fig. 17 is a graphical representation of additional aspects of the present
invention;
Fig. 18 is a graphical representation of additional aspects of the present
invention; and
Fig. 19 is a graphical representation of still further aspects of the present
invention.
DETAILED DESCRIPTION
In the following detailed description, reference is made to the accompanying
drawings
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that form a part hereof, and in which is shown by way of illustration,
specific embodiments in
which the invention may be practiced. These embodiments are described in
sufficient detail to
enable those skilled in the art to practice the invention, and it is to be
understood that other
embodiments may be utilized. It is also to be understood that structural,
procedural and
system changes may be made without departing from the spirit and scope of the
present
invention. In addition, well-known structures, circuits and techniques have
not been shown in
detail in order not to obscure the understanding of this description. The
following detailed
description is, therefore, not to be taken in a limiting sense, and the scope
of the present
invention is defined by the appended claims and their equivalents.
As will now be described in detail, embodiments of the present invention
relate to
upgrading a continuously flowing stream including heavy crude oils, extra-
heavy crude oils,
ultra-heavy crude oils, bitumen, and the like without limitation in regard to
exact composition or
geologic or geographic origin, which hereinafter are referred to inclusively
as heavy crude oils or
simply HCO. Indeed, as used herein, the term "heavy crude oil" and/or "HCO"
refers to
substantially any crude oil or hydrocarbon-containing material measuring at or
below about 22.3
API, with lower API values corresponding to higher densities. Referring to
Fig. 1, in one
example, a system 20 is provided for the hydrothermal upgrading of a process
stream (reaction
mixture) 22 (HCO and water) by a reactor portion (section) 24 configured to
progressively heat
the process stream 22 as a function of the reaction coordinate (R.C.) 26. The
R.C. 26 may be
calculated as (a) the relative distance between an inlet 28 and outlet 30
traversed by the process
stream 22 within the reactor section 24, or (b) the time elapsed after the
reaction mixture 22
enters the reactor section at 28, times flow rate divided by the total volume
of the fluid flow path
within the reactor section 24. In Fig. 1, the rate of heat applied (rate of
heat transfer or thermal
flux) to the reaction mixture 22 is shown graphically as temperature (T) of
the process mixture
22 (on the y-axis), as a function of the R.C. 26 (on the x-axis.) Moreover, in
particular
embodiments, the thermal flux or temperature applied at the reactor 24 is
increased progressively
along the downstream direction a. This progressively increasing thermal flux
may be provided,
for example, by an otherwise conventional counter-flow heat exchanger such as
the shell-and-
tube heat exchanger shown and described hereinbelow with respect to Figs. 7, 8
and 11. It should
be recognized, however, that substantially any type of heater known to those
skilled in the art,
e.g., arranged in series with one another along the process fluid flow path
may be used, without
departing from the scope of the present invention.
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In particular embodiments, the reactor section 24 includes one or more process
flow
tubes each having an interior cross-sectional dimension (e.g., diameter) in a
plane extending
transversely to the downstream direction a therethrough. In this regard, it
should be
recognized that the process flow tubes may be disposed in series, such as
shown and
described hereinbelow with respect to Fig. 7, and/or in parallel, as shown in
Fig. 11.
Regardless of whether the flow tubes are disposed in series, in parallel, or
in a combination
thereof, the flow tubes are provided with a combined length of at least about
30 times the
aggregated interior cross-sectional dimension. For purposes of computing the
aggregated
interior cross-sectional dimension, flow tubes disposed in series are treated
as a single tube.
Therefore, for example, a reactor having one or more flow tubes in series,
each with a
diameter of 5 cm would have an aggregated cross-sectional dimension of 5 cm
and a total
length of at least 30 x 5 cm or about 150 cm. Similarly, a reactor having
three parallel flow
tubes each having a diameter of 5 cm, would have an aggregated cross-sectional
dimension of
cm and a length of at least about 450 cm. Moreover, although these examples
contemplate
15 flow tubes of circular cross-section, one skilled in the art will
recognize that tubes of
substantially any shape cross-section, such as square, oblong, etc., may be
used without
departing from the scope of the present invention.
The reactor 24 is configured to apply heat to the reaction mixture flowing
therethrough, to progressively heat the reaction mixture 22 so that the
reaction mixture is
disposed at a lower temperature at an upstream or inlet portion of the
reactor, e.g., at 28, than
at a downstream or outlet portion of the reactor, e.g., at 30. In particular
examples, the reactor
24 is configured to progressively heat the reaction mixture 22 from an inlet
28 temperature of
about 60 C to 200 C, to an outlet 30 temperature (T(max)1) of between about
260 C and
400 C. It should be noted that this progressive heating may be accomplished
either
substantially continuously, as shown in Figs. 1-7, or discontinuously, as will
be discussed in
greater detail hereinbelow with respect to Figs. 8-10. It is also noted that
the reactor 24 is
configured to maintain the reaction mixture 22 at a pressure sufficient to
ensure that the
reaction mixture remains a single phase at T(max)1, i.e., to substantially
prevent formation of
a gas phase separate from the liquid phase of the reaction mixture 22. In
various exemplary
embodiments, pressure within reactor 24 may be maintained within a range of
about 1500 to
about 3000 psia, with particular embodiments being maintained within a range
of 1500 to
2000 psia, and other embodiments being maintained within a range of about 2000
to 3000
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psia.
As also shown, system 20 may also include a controller 32, e.g., in the form
of a
conventional closed-loop programmable logic controller (PLC) or process
automation
controller (PAC) such as the model T2750 commercially available from Foxboro
(Invensys
Systems, Inc., Foxborough, MA, USA), optionally augmented with model
predictive control
(MPC) capability, communicably coupled to reactor 24, including a flowmeter
and
temperature and pressure probes associated therewith (not shown) for capturing
the flow rate,
temperature and pressure of the process stream 22. The controller 32 is
configured to adjust
both the rate of flow of reaction mixture 22 into the reactor 24, and/or the
rate of heat applied
(rate of heat transfer or thermal flux) to the reaction mixture 22 in the
reactor 24 (e.g., by
controlling operation of hardware commonly associated with process flow, such
as pumps,
valves, heaters, etc. (not shown)). In particular embodiments, controller 32
is configured to
ensure that the flow rate is sufficiently high and the rate of heat transfer
is sufficiently low to
minimize or substantially prevent coke formation, while maintaining a total
residence time of
.. the reaction mixture 22 within the reactor 24 of greater than about 1
minute and less than
about 25 minutes, calculated as the total volume of the fluid flow path within
the reactor
divided by the flow rate, to form a product mixture 34 exiting the reactor at
30. It is noted that
both the flow rate and the thermal flux from the inside surface of the tube or
tubes in the
reactor section may be optimized to minimize or prevent coke formation while
achieving the
desired level of upgrading and maximizing throughput, while taking into
consideration the
thennal conductivity of the reaction mixture 22.
As also shown, system 20 includes a recovery portion (section) 38 configured
to receive
the process stream, which has now been transformed into product mixture 34,
exiting the reactor
section at 30. Recovery section 38 is configured to reduce the temperature of
the product mixture
34, e.g., to between 60 C and 200 C, and to also effect a corresponding
reduction in the vapor
pressure of the mixture 34. It is also noted that in particular embodiments,
recovery section 38
includes a water separator 40 configured to separate water from the upgraded
crude oil, which
exit the recovery section 38 at 42 and 44, respectively.
An aspect of the present invention is thus the gradual heating of reaction
mixture 22,
including an HCO stream and water, flowing through reactor section 24, on a
time scale
configured to promote, at relatively low temperatures, the disaggregation of
HCO components
and their substantially uniform distribution in the matrix of the reaction
mixture, and additionally
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at higher temperatures upgrading reactions, all the while minimizing or
preventing coking. It is
noted that the use of tube reactors in petrochemical processing to effect
chemical transformations
is commonplace. For example, it is the standard approach for cracking of gas-
phase naphtha at
temperatures in excess of 800 C to produce ethylene. It is noted, however,
that the use of tube
reactors to effect the chemical transformation of liquids at the relatively
lower temperatures is
rather uncommon or absent in commercially relevant processes used in either in
refining or
petrochemical operations.
Another aspect of the present invention is that instead of the reactor having
a
substantially uniform temperature distribution, the flowing mixture 22 instead
experiences a
deliberately non-uniform application of heat (thermal flux) between the inlet
28 and the outlet
30. Though not wishing to be bound by any particular theory of operation, the
belief is that the
aforementioned approach facilitates upgrading by fostering sequentially two
different physico-
chemical processes. First, as discussed briefly above, the use of time and the
application of
progressively increasing temperatures between the inlet and outlet of the
reactor section serves to
disintegrate physical structures in HCO and/or effect the dissolution of HCO
components to
yield a substantially uniform dispersion by the point where the mixture
reaches a temperature of
about 80% to about 90% of the predetermined maximum temperature at some point
before the
outlet. The process of disaggregation, disintegration, or destructuring of
assemblages of HCO
components and the dispersing and/or dissolution of the same will be
inclusively referred to
hereinafter as the disaggregation reaction or simply disaggregation. Through
disaggregation,
asphaltenes and other heavy compounds that are generally associated in HCO are
thought to be
dispersed and nominally separated from one other, predisposing them to undergo
upgrading
reactions involving water and minimizing the possibility that they thereafter
will undergo
retrograde reactions with each other that lead to the formation of more and
larger asphaltenes
molecules and possibly coke. The process that yields product qualities such as
density and
viscosity that are improved over those of the HCO feed due to upgrading
reactions involving
heavy components originating in HCO will be referred to inclusively as the
upgrading reaction or
simply upgrading.
It will be understood that the embodiments shown and described herein for
upgrading
HCO do not purport to selectively and/or exclusively promote first the initial
disaggregation
reaction and subsequently the upgrading reaction, nor is there a presumption
that the latter occurs
only when the temperature of the reaction mixture reaches and exceeds
temperatures of between
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about 80% and 90% of the predetermined maximum temperature. Rather, the
reaction mixture
will predominantly undergo disaggregation reactions at lower temperatures in
the reactor section
while upgrading reactions occur predominantly at the higher temperatures in
the reactor section.
Indeed, as shown in Fig. 10 (discussed in greater detail hereinbelow) an
aspect of embodiments
of the present invention is the fostering of the disaggregation of HCO
components prior to their
being subjected to conditions of elevated temperature at which upgrading
reactions occur,
maximizing the efficiency and extent of upgrading at the highest temperatures
while minimizing
undesirable side reactions that lead to coke formation.
Turning now to Fig. 2, the controlled, progressive increase in the temperature
of the
reaction mixture between the reactor section inlet and outlet is but one
aspect of the present
invention that preferentially promotes, first, the disaggregation reaction and
then the upgrading
reaction. An additional approach for promoting disaggregation and upgrading
reactions
associated with various embodiments of the present invention involves the
selection and
contacting of the HCO stream in a premix section with materials selected to
promote one or both
.. of those reactions. In one embodiment of the invention, HCO flowing through
a premix section
50 is contacted with a material including either water or steam at a
temperature at or below the
desired predetermined inlet 28 temperature of the process stream 22, e.g., at
a temperature at or
below about 200 C. In another embodiment the temperature of the water or
steam contacting
HCO flowing through the premix section may be as high as about T(max)1 or
about 350 C,
whichever is lower so as to avoid the promotion of localized cracking of HCO
components at or
near the point of contacting, which is thought to lead to coke formation. The
mixture of HCO
and this water or steam becomes the process stream 22 that is fed to the
reactor section 24 at inlet
28.
Not wishing to be bound by any particular theory of operation, it is believed
that the
contacting of HCO, which has not substantially undergone disaggregation, with
water whose
temperature is greater than about 325 C may promote coking due to localized
high rates of
cracking at or near the point of contacting followed by retrograde
intermolecular reactions of
components that are not substantially disaggregated within the reaction
mixture 22. Thus, while
coke formation by this mechanism may be minimized by ensuring that HCO
components are
substantially disaggregated prior to contacting with water that is
supercritical (temperature and
pressure are equal to or greater than about 374 C and 3200 psia) or near-
supercritical (e.g.
temperature and pressure are in the range of about 325 C to 374 C and 2000
psia to 3200 psia,
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respectively), it is expected to be reduced even further through the promoting
of disaggregation
and the contacting with water whose temperature is less than about 325 C.
Thus, as shown, system 220 of Fig. 2 is substantially similar to system 20 of
Fig. 1, while
also including an optional premix section 50 for contacting the HCO with water
or steam to form
the process stream 22. As also shown, the recovery section of system 220
includes an optional
energy recovery subsection (e.g., heat exchanger) 52, which is configured to
recover thermal
energy from the product mixture 34 and to distribute the recovered energy to
the reactor section
24 as shown at 56. The energy removed from the mixture 34 is shown graphically
as a reduction
in temperature (T) as a function of R.C. 26. Still further, system 220 may
include an optional
water recycling loop 58 configured to recirculate the water 42 recovered at
water separator 40, to
the premix section 50, although it will be understood that other embodiments
water used in
contacting HCO in the premix section and/or the reaction mixture in the
reactor section
(discussed in greater detail hereinbelow) may be from sources instead of or in
addition to water
from the recycling loop.
Turning now to Fig. 3, in another variation of the foregoing, a system 320 is
substantially
similar to system 220, with the addition to the recovery section of an
optional light hydrocarbon
removal device 62. An example of a suitable device 62 may include a
conventional flash drum
configured for recovering light hydrocarbons (e.g., naphtha, distillates,
condensates and the like,
hereinafter referred to simply as LHC) from the product mixture 34. The
recovered LHC may
.. then be recirculated via hydrocarbon recycling loop 64 back to the premix
section 50, to help
promote the disaggregation reaction.
Turning now to Fig. 4, in yet another variation of the foregoing embodiments,
a system
420 is substantially similar to system 320, with the optional injection of
water or steam (e.g.,
from recycling loop 58) at one or more points in reactor section 24 instead of
the premix section
50. This effectively provides for contacting the HCO stream with the
hydrocarbons, and
therefore promoting the disaggregation reaction, prior to contacting the HCO
with water or
steam, which may be particularly effective for predisposing the reaction
mixture 22 toward
mixing with injected water or steam and undergoing upgrading reactions
involving water at the
higher temperatures found in the reactor section 24.
In another variation shown in Fig. 5, a system 520 is substantially similar to
system 420,
with the optional injection of light hydrocarbon from recycling loop 64 into
the reactor section
24 as well as premix section 50. It is noted that in particular embodiments,
the injection of light
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hydrocarbon occurs at one or more points 68 prior to where the reaction
mixture 22 has reached
a temperature T85 of between about 80% to 90% of the temperature T(max)1. In
other
embodiments, the injection of light hydrocarbon occurs at one or more points
68 prior to where
the reaction mixture 22 has reached a temperature T65 of between about 60% to
70% of the
temperature T(max)1.
Turning now to Fig. 6A, system 620 is substantially similar to system 520,
with the
injection of light hydrocarbon and water/steam into reactor 24 without any
injection into premix
section 50. Thus, in this embodiment, no material is mixed with the HCO in the
premix section
(the premix section 50 is thus not required, although shown), although the HCO
may contain
water and/or LHC as the result of the steam assisted gravity drainage (SAGD)
process in
widespread use in Alberta for extracting and producing HCO from oil sands
deposits, which
HCO may be upgraded by embodiments of the present invention. Rather, water or
steam are
injected at one or more points in the reactor section 24 prior to the T85 or
T65 points as discussed
hereinabove. In this embodiment, it is also noted that at the point of
injection the temperatures of
the water or steam supplied by loop 58 may be equal to or less than about 80%
to 90% of the
predetermined maximum temperature in the reactor section T(max)1. In another
embodiment,
the temperature of the water injected into the reactor section may have
temperatures greater than
T(max)1 but less than about 350 C. System 620' of Fig. 6B is substantially
similar to system
620, but with the additional injection of water/steam and hydrocarbon via
loops 58 and 64,
respectively, into premix section 50.
It is noted that in any of the embodiments shown and described herein, the
contacting of
water/steam and/or hydrocarbon with the HCO may be facilitated by a variety of
conventional
means including but not limited to mechanical stirring, inline mixing, static
mixing, a mixing
eductor, a radial (vortex) premixer, and/or a pump that continuously drives
the reaction mixture
22 from the premix section 50 into the reactor section 24. Moreover, the
examples shown and
described herein are not intended to be limiting, with other combinations of
injecting
water/steam and hydrocarbons into the premix and reactor sections being
included within the
scope of the present invention.
Still further, in various embodiments described herein, the amount of
water/steam and
hydrocarbon supplied to the HCO in process stream 22 is configured to provide
a final ratio of
HCO to water (HCO:water) ranges from about 1:1 and 20:1, while the ratio of
native HCO to
LHC not native to the HCO (HCO:LHC) ranges from about 1:2 and 20:1. It is
noted that the non-
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native LHC may be present in the HCO stream flowing into the premix section 50
or may be that
which is introduced in either the premix section 50 or the reactor section 24.
The selection of the particular final value for HCO:water may be based on the
balancing
of two opposing factors. Higher relative concentrations of water may be
beneficial in that they
foster improved heat transfer from the walls of tube in the reactor section
24, suppress coke
formation by quenching or preventing reactions between HCO components, and
improve the
kinetics of upgrading reactions involving water. Some disadvantages of
relatively high water
concentrations relate to the fact that displacement of HCO by water reduces
the effective
throughput of HCO while increasing operating costs due to the need to invest
thermal energy to
heat not only HCO but also water. Given that the heat capacity of the latter
is approximately
twice that of HCO and other hydrocarbons, each incremental increase in the
water content of the
reaction mixture requires proportionately more thermal energy to heat the
reaction mixture to
T(max)1. Because, in comparison with water, added light hydrocarbon is thought
to be
particularly effective for promoting the disaggregation reaction while
suppressing the formation
of coke, particular embodiments of the present invention, as discussed
hereinabove, involve the
initial contacting of the HCO by light hydrocarbon in the premix section 50
and the injection at
one or more points later in the process (e.g., in the reactor section 24) of a
minimum amount of
water required to effect the desired level of upgrading without formation of
unacceptable levels
of coke. It will be understood that examples given hereinabove, which depict
the contacting of
HCO in the premix section and/or the reaction mixture in the reactor section
by LHC recovered
directly from the product mixture by the LHC removal device 62, are
nonlimiting. In other
embodiments, LHC may be used for such purpose which come from sources other
than or in
addition to the optional device 62.
Referring now to Fig. 7, in a nonlimiting example, the reactor section 24 is a
conventional shell-and-tube heat exchanger in which the reaction mixture 22
flowing through a
central tube is heated by a heating fluid 70 flowing in the direction opposite
that of the reaction
mixture 22. The heating fluid 70 flowing into the shell has a temperature
sufficient to ensure that
the temperature of the reaction mixture 22 at the outlet 30 of the reactor
section 24 is at about the
predetermined maximum temperature T(max)1. The flow rate and temperature of
the heating
fluid are adjusted to create a continuously-varying temperature profile along
the length of the
reactor section 24 as shown and described hereinabove with respect to the
graphical components
of the Figures. The temperatures of the reaction mixture 22 and the heating
fluid 70 are shown
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graphically as a function of the reaction coordinate 26. The graphical
component of Fig. 7 also
indicates that the occurrence of the disaggregation reactions predominates at
the inlet/low
temperature end 28 of the reaction section 24, while the upgrading reactions
predominate at the
outlet/high temperature end 30 of the reaction section 24.
As also mentioned hereinabove, although reactor section 24 may take the form
of a
single-tube heat exchanger, the skilled artisan will recognize that reactor
section 24 may
alternatively include a heat exchanger having a plurality of parallel tubes
within the shell,
wherein the inlets of all the tubes are communicably coupled by a common inlet
chamber and the
outlets are communicably coupled by a common outlet chamber, such as shown in
Fig. 11.
Turning now to Fig. 8, an alternate reactor section shown at 24' may take the
form of a
series of shell-and-tube subsections 72, where T(hf); is the temperature of
the heating fluid 70 at
the inlet to each subsection. Reactor section 24 is otherwise substantially
similar to reactor
section 24, such as described with respect to Fig. 7.
Referring to Fig. 9, an optional reactor section 24" is substantially similar
to reactor
section 24', but with the interposition of thermal-soak chambers 74 disposed
serially between the
subsections 72. As shown graphically, the thermal-soak chambers 74 are
configured as insulated,
unheated flow-through chambers that effectively lengthen the residence time of
the reaction
mixture 22 at various temperatures as the mixture 22 flows through the reactor
section 24". As
shown graphically, the thermal-soak chambers 74 effectively provide a
substantially step-wise
increase in temperature as a function of the R.C. 26, e.g., along the length
of the reactor section
24". The particular temperatures and the resulting residence time at those
temperatures may be
selected to facilitate kinetics related to the disaggregation and upgrading
reactions. Moreover,
although in particular embodiments the thermal-soak chambers 74 will be
unheated, it should be
recognized that in particular applications, it may be appropriate to apply
some amount of heat to
the thermal-soak chambers, such as may be desired to maintain the reaction
mixture flowing
therethrough at an approximately uniform temperature.
As discussed above, the graphical component indicates that the occurrence of
the
disaggregation reactions predominates at the inlet/low temperature end 28 of
the reaction section
24", while the upgrading reactions predominate at the outlet/high temperature
end 30 of the
reaction section 24". It should be recognized that any number of the reactor
subsections 72 of
Figs. 8 and 9 may be used, depending on the particular application. It should
also be recognized
that in various embodiments, pressure within the various components of
reactors 24, 24', 24", is
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maintained at levels sufficient to prevent the formation of a phase separate
from the liquid phase
of reaction mixture 22, as discussed hereinabove.
Turning now to Fig. 10, any of the aforementioned reactor portions 24, 24',
24" may be
further modified to include a supplemental reactor section 80 serially
disposed at outlet 30
thereof. It is noted that supplemental reactor section 80 may be substantially
similar to reactor
section 24, 24', 24" and/or one or more reactor subsections 72. As shown,
reactor section 24,
24', 24" effectively brings the reaction mixture 22 to T(max)1 as described
hereinabove, while
the supplemental reactor section 80 is configured so that the reaction mixture
flowing
therethrough achieves a predetermined maximum temperature T(max)2 . In
particular
embodiments, T(max)2 is within a range of approximately 1.0 to 1.1 times
T(max)1, and in other
embodiments, is within a range of approximately 1.1 to 1.4 times T(max)1, as
will be discussed
in greater detail hereinbelow. Examples of reactor sections 24, 24', 24" and
energy recovery
sections 52 which include multiple parallel process flow paths as discussed
hereinabove,
disposed in series, are shown in Fig. 11.
In the foregoing embodiments employing serial heat exchangers, the temperature
and rate
of heating fluid flow through each shell may be individually controlled to
control the temperature
and rate of heat applied to the reactions mixture 22. Moreover, as also
discussed hereinabove, the
reaction mixture is maintained at pressure sufficient to maintain the reaction
mixture as a single
phase throughout the reactor section. Doing so is important from the viewpoint
of process
energetics because in general, liquid-to-gas phase changes consume significant
energy as a
function of the heat of vaporization, AHvap . The importance of this issue is
particularly acute
given that AHvap for water may be 5-9 times higher than for many lower
molecular weight
hydrocarbons. Therefore, in a particular embodiment of the present invention,
the reactor section
is operated at pressures in excess of those required to maintain the reaction
mixture in liquid
phase when the reaction mixture experiences the maximum temperature in the
reactor section.
Higher pressures may also enhance the kinetics of the disaggregation reaction,
the upgrading
reaction, or both. Yet, any such benefit may be offset by higher equipment
costs and operating
costs, which include equipment maintenance. Therefore, in a particular
embodiment of the
present invention the reactor section is maintained at an operating pressure
that is approximately
5% to 10% in excess of that required to maintain the reaction mixture as a
liquid phase. That
pressure varies as a function of the predetermined maximum temperature in the
reactor section;
the amount and chemical composition of LHC in the reaction mixture, whether
they are native to
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the HCO, added to enhance disaggregation reactions, or generated through
upgrading reactions;
and the amount of water in the reaction mixture. Various embodiments use
pressures within a
range of from about 1500 to about 3000 psia. Under some conditions of
temperature and reaction
mixture composition, the required operating pressure may be approximately 2000
psia, while
under other conditions the required pressure may be from about 1500 psia to
about 2000 psi,
while under yet other conditions the required pressure may be from about 2000
psi to
approximately 3000 psia.
In various embodiments of the present invention, the predetermined maximum
temperature at the outlet from the reactor section is kept as low as possible
in consideration of
energy costs, the aforementioned costs associated with building and
maintaining processing
equipment, and of the desire to minimize coke formation. However, doing so
works at
cross-purposes to the promotion of the disaggregation and upgrading reactions,
which are
enhanced as a function of increasing temperature. Consequently, another
processing variable that
plays a role in chemical kinetics must be exploited, namely, time.
The shell-and-tube configuration of the reactor section provides two important
benefits,
one stated and one implied. The former relates to the possibility for creating
a deliberately
non-isothermal temperature regime to preferentially promote disaggregation
before applying
maximum thermal energy to achieve predetermined maximum temperatures and drive
upgrading
reactions involving water. The implied benefit is the well-known enhancement
of heat transfer
by means of the high ratio of surface area to volume (surface:volume)
available in tubular
reactors. Yet, the latter benefit is achieved at the price of pressure drop
between the inlet and
outlet as a function of increasing tube length and decreasing tube cross-
sectional dimension, both
of which increase surface:volume. Therefore, the idea of increasing the total
time spent by the
reaction mixture in the reactor section by increasing the tubing length seems
dubious. And
though the problem may be mitigated by a variety of means, it ultimately is
bounded by the
interplay between three variables: the number and lengths of the tubes inside
of a shell-and-tube
system; the viscosity of the mixture flowing through them; and flow rate. The
possibilities for
reducing viscosity by increasing temperature or decreasing the amount of HCO
relative to added
light hydrocarbon or water are limited, as these measures tend to work in
opposition to other
considerations related to upgrading optimization. Similarly, reducing the flow
rate serves
primarily to reduce the throughput of HCO.
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The present inventor's solution to the problem resides in an approach that
reduces
pressure drop by effectively increasing the diameters of tubes in the reactor
section, which at first
seems counterintuitive and even contrary to a key benefit noted for tubular
reactors, namely, that
they offer a high ratio of surface area to volume. However, in particular
embodiments of the
present invention, one or more flow-through chambers (thermal soak chambers)
74, that are only
insulated and not heated will be interposed between individual reactor
subsections, where the
volume of the flow-through chambers 74 is equal to about the desired increase
in residence time
multiplied by the flow rate through the reactor section. For example, a single
such chamber 74 of
volume Vc located in the reactor section before the T85 point could increase
by an average of one
minute the time for the disaggregation reaction if Vc = 1 minute x FRm, where
Fpm = the flow
rate of the reaction mixture through the reactor section in units of
volume/minute. (This does not
take into account, of course, the consequence of flow-based mixing that might
occur in the flow-
through chamber.) Similarly, in another particular embodiment of the
invention, the reactor
section 24, 24', 24" is extended by the location of an insulated, unheated
flow-through chamber
74 of volume Vc = m x FRm at the outlet of the last shell-and-tube subsection
in the reactor
section where the reaction mixture is heated to about the predetermined
maximum temperature
T(max)1, resulting in the increase by m minutes the time available for
upgrading reactions,
where FRm is the flow rate of the reaction mixture in units of volume/minute.
In yet another
particular embodiment, the throughput of the upgrading system may be increased
without
significantly increasing the pressure drop across the reactor section by
increasing the size of the
flow-through chamber 74 installed at the outlet from the last shell-and-tube
subsection, the size
and number of flow-through chambers 74 interposed between the inlet and
outlet, and the
number of tubes contained in the shell-and-tube subsections 72. More than one
shell-and-tube
subsection 72 may be installed between the chambers 74.
Though not wishing to be bound by any particular theoretical reasoning,
upgrading
reactions in the reactor section of the instant invention are thought to be of
two general types
discussed briefly hereinabove. In the heterolytic scission of the covalent
bond between two
atoms, the electron pair is divided asymmetrically whereas homolytic scission
results in the
bonding electron pair being divided equally between the two bonded atoms. For
convenience, the
heterolytic and homolytic reactions of HCO components are discussed herein
will be referred to
hereinafter as upgrading reactions Type I and Type II, respectively. In the
absence of catalysts,
homolytic reactions may be promoted at elevated temperatures, an important
example being the
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production of ethylene through the gas-phase cracking of naphtha at
temperatures of about 850
C. Being highly endothermic (AH has a high positive value), the entropy term
of the Gibbs
Free Energy equation AG = AH ¨ TAS prevails under such conditions. By
contrast, heterolytic
reactions may be promoted by catalysts that facilitate reactions by lowering
the activation
energy, which is generally desirable as this allows reactions to occur at
lower temperatures than
would be possible without a catalyst. An example is fluidized catalytic
cracking (FCC) in
refineries where FCC units play a vitally important role by increasing the
yields of gasoline
obtained from crude oil through the cracking of heavier hydrocarbons to form
lighter ones. FCC
catalysts are fine powders that function as a substrate onto which
hydrocarbons adsorb in order
for catalysis to occur. Other catalysts may be molecules that promote
reactions by participating
in them, in some cases being chemically integrated into reaction
intermediates, but always being
regenerated. The ability of water to function in this way is well known, and
in a particular
embodiment of the present invention, the reaction mixture in the reactor
section is heated to a
final predetermined temperature T(max)1 of about 260 C to about 325 C (or
about 260 C to
about 400 C in some embodiments) such that the upgrading rates due to water-
catalyzed Type I
reactions are thought to become significant when the reaction mixture reaches
a temperature of
about 80% to about 90% of T(max)1, where significant levels of homolytic
cracking do not
occur.
Though not wishing to be bound by any particular theory of operation, in
particular
embodiments of the invention the reaction mixture flows through a primary
reactor section 24,
24', 24" and is heated to the predetermined maximum temperature T(max)1 where
water-catalyzed Type I reactions are thought to occur without Type II
reactions occurring at an
appreciable rate, and then flows into a supplemental reactor section 80 (Fig.
10) appended to the
primary reactor section where heating of the reaction mixture continues to a
second
predetermined maximum temperature of T(max)2. In this particular embodiment,
the flow rate
of, and the thermal flux into the reaction mixture are controlled to minimize
or prevent coke
formation in the reaction mixture as it is heated progressively while flowing
through both the
primary and the supplemental reactor sections.
In a particular embodiment, the predetermined maximum temperature T(max)2 of
the
supplemental reactor section 80 is between about 1.0 and 1.1 times T(max)1,
the principal
purpose of the supplemental reactor section being in this case to further
drive Type I reactions to
an extent beyond that which might have been achieved in the primary reactor
section without
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appreciably fostering Type II reactions. In yet another embodiment, the
supplemental reactor
section raises the reaction mixture to a predetermined maximum temperature
T(max)2 of
between about 1.1 and 1.4 times T(max)1 to promote further upgrading by means
of Type II
reactions in addition to Type I reactions. Though not wishing to be limited by
any particular
theoretical reasoning, it is suspected that the benefit of such a strategy is
that promoting first the
Type I upgrading reactions in the primary reactor section reduces the average
molecular weight
of components in the reaction mixture to produce lighter compounds, which in
turn serve to
promote further disaggregation and even the dissolution of heavy components,
minimizing the
potential for coke formation or the retrograde formation of asphaltenes and
other heavy
components when subjected to thermal conditions conducive to Type II upgrading
reactions in
the supplemental reactor section.
In particular embodiments as discussed hereinabove, the recovery section 38
reduces the
temperature of the product mixture to between about 60 C and 200 C. The
reduction in
temperature results in a corresponding reduction in the vapor pressure of the
product mixture and
an increase in viscosity. Moreover, the product mixture 34 (Fig. 3) may flow
under pressure
through water separator 40, which may take the form of one or more liquid-
liquid separators for
the purpose of removing water. As discussed, recovered water can then be
recycled in the
process at 58 or used elsewhere, for example, it may be combined with water
that is converted to
steam for downhole injection in the context of the SAGD process for extracting
HCO form oil
sands. And as also discussed, in various embodiments, the product mixture from
which water has
been substantially removed flows into flash drum 62, having had its
temperature reduced to a
level appropriate to remove desired quantities of LHC at 64, e.g., by flash
distillation of vapors,
which LHC are referred to as condensate. The flash drum 62 may be
appropriately configured to
recover other volatile components in the product mixture that are lighter than
condensate, which
may include methane ethane, propane, hydrogen sulfide, CO2, COS, CS2, SO,, and
possibly
nitrogenous compounds. In various embodiments as discussed above, the
recovered condensate
is used to facilitate the disaggregation reaction by contacting with HCO in
the premix section 50
or injection into the reactor section 24, 24', 24".
Scalability is another among the many aspects of the present invention that
have already
been identified, and which are believed to be novel, non-obvious and
beneficial. In particular,
embodiments of the invention may be scaled without substantially changing the
basic design to
permit the construction of upgrading systems with capacities ranging from <
lbarrel per day
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(bpd) to >10,000 bpd. This feature means that the present invention combines
the best features of
the two upgrading methodologies that enjoy broad use in Canada today. First,
like the large,
expensive technology used to produce synthetic crude oil, it offers the
possibility to upgrade
HCO to meet pipeline specifications without relying on dilution. Second, like
diluent-based
upgrading, the system and method described herein is suitable for distributed
implementation on
a smaller scale at or near wellheads and remote production fields.
Another aspect of the present invention is that particular embodiments
maximize energy
efficiency and ecofriendliness. For example, the recovery section 38 captures
and recycles
thermal energy and also water for reuse in the system and method described
herein for upgrading
HCO. Particularly beneficial is the ability to substantially minimize the
thermal energy required
to heat the reaction mixture to T(max)1, or to T(max)2, by substantially
maximizing the ratio
HCO:water. When the reaction mixture comprises only HCO and water, relatively
large amounts
of the latter are required because, apart from being involved in upgrading
reactions, it is thought
to be important for facilitating the dissociation and dispersal of HCO
components in a
water-dominated matrix. This latter function would not merely expose
individual molecules such
as asphaltenes to the upgrading reactions, but also suppress counterproductive
retrograde or
intramolecular reactions that could lead to the formation of undesirable
products such as coke.
By this strategy, the required amount of water relative to HCO is much greater
than the
minimum required only for actual upgrading reactions, and an energy penalty
attaches to it
because the heat capacity of water is approximately twice that of HCO.
Embodiments of the
instant invention address this dilemma by the mixing of HCO in the premix 50
and/or reactor 24,
24', 24" sections with LHC, e.g. those optionally recovered from the product
mixture and
recycled to the front end of the process, which are thought to facilitate
disaggregation reactions
without themselves being subject to appreciable additional chemical
transformation in the reactor
section. Subsequently, only the amount of water required to foster water-based
upgrading
reactions is added to the reaction mixture flowing through the reactor
section. Not only does this
approach help minimize energy requirements, but it offers the possibility to
improve the kinetics
of the disaggregation reaction because the LHC are presumably more effective
than water in that
regard.
Additional aspects of the invention concern the specific, deliberate design of
the reactor
section(s) to effect the gradual heating of the reaction mixture for the
explicit purpose of a)
avoiding levels of thermal flux into reaction mixture from the inside walls of
the reactor to
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minimize or prevent coking; b) promoting disaggregation reactions by means of
optionally
contacting of HCO with LHC in the premix section and/or the reactor section;
c) promoting
disaggregation reactions prior to the reaction mixture reaching a temperature
slightly below (e.g.,
between about 80% and about 90% of) the predetermined maximum temperature
T(max)1; d)
promoting upgrading reactions in a reactor section at temperatures between
that point (e.g., about
80% to 90% of T(max)1) and about T(max)1 where Type I upgrading reactions are
thought to
occur; e) avoiding subjecting the HCO to localized heating at the point of
injection of and
contacting by water that has been heated to supercritical or near-
supercritical temperatures when
the HCO has not already undergone disaggregation, under which conditions the
localized rate of
heating exceeds the rate of disaggregation and instead promotes the formation
of coke due to
intermolecular reactions between HCO components and/or localized high rates of
cracking,
ostensibly by homolytic mechanisms; and f) optionally promoting further
upgrading reactions in
the reaction mixture flowing through a supplemental reaction section whose
predetermined
maximum temperature T(max)2 is between (i) 1.0 and 1.1 times T(max)1 where
Type I reactions
are thought to predominate, or (ii) 1.1 to 1.4 times T(max)1 where Type H
upgrading reactions
are thought to occur at the point in the supplemental reactor when the
reaction mixture reaches a
temperature greater than about 1.1 times T(max)1. It should be noted the
various means
described by which the invention minimizes the formation of coke effectively
increase the yield
of upgrading because coking converts HCO components into a low-value byproduct
that has
little value other than as fuel.
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Heretofore the underlying concept of the instant invention has been to
introduce the HCO
feed to an upgrading process that is more or less linear: the feed flows
continuously into a reactor
or a series of reactor subsections after optionally flowing through a premix
section; a reaction
mixture is generated by contacting the HCO with LHC and/or water/steam at one
or a plurality of
points in the premix and/or reactor sections; the mixture is progressively
heated in the reactor or
plurality of reactor subsections to a predetermined maximum temperature
T(max)1; it optionally
flows through a supplemental reactor section where it is heated progressively
to a second
predetermined maximum temperature T(max)2; and the product mixture flowing
therefrom is
processed through a recovery section designed to remove water from the
upgraded HCO product,
recover thermal energy, and optionally recover LHC to be recycled to one or
more upstream
points for contacting with the HCO to aid in disaggregation believed to be an
important for
predisposing HCO to upgrading reactions with water. (The foregoing summary is
provided by
way of illustration, is not intended to describe all possible embodiments, and
therefore is not
limiting.) However, in other embodiments of the invention, one or more
additional elements may
be added into the process flowpath to create one or more branch points in the
flowpath. These
embodiments, which will now be described in greater detail, are intended to
increase yet further
the selectivity, efficiency, and effectiveness of the upgrading process. This
in turn tends to
increase control over the range and quality of products that can be obtained
from the process,
thereby opening up distinctive, new possibilities to enhance the economics of
producing HCO by
the SAGD process and upgrading it for transport to market.
Referring to Figure 12, the reaction mixture flows through the reactor section
24, 24',
24" such as shown and described hereinabove, including one or more subsections
and optional
thermal soak chambers interposed in the flow path. The product mixture then
flows from the
outlet 30 of the primary reactor section 24, etc., into a flash distillation
column 90, also
designated FDC1, wherein the product mixture is separated into two product
fractions
differentiated on the basis of the boiling point ranges for their respective
components, which for
simplicity will be referred to as light and heavy fractions 92 and 94,
respectively. The light
fraction 92 includes water and light hydrocarbons (LHC), the latter being
designated LHC1 and
having boiling points at atmospheric pressure below a predetermined maximum
value of
BP(LHC1), which in one embodiment is below about 280 C, while in yet another
particular
embodiment BP(LHC1) may be below about 220 C, and in yet another particular
embodiment,
it may be less than about 160 C. The heavy fraction 94, which may be referred
to as the
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predominantly heavy hydrocarbon (HHC) product stream, the HHC residue, or
simply HHC1,
includes components in the product mixture 34 whose boiling points are higher
than BP(LHC1).
Components in light fraction 92 from FDC1 90 are cooled in a condenser (not
shown) to yield
liquid mixture that is separated into an aqueous phase and a hydrocarbon
phase. By contrast, the
HHC1 94, flows without having been cooled from the bottom of the distillation
column 90 into
an optional premix section 50' configured in the flow path between FDC1 90 and
a supplemental
reactor section 80 such as shown and described hereinabove.
Not wishing to be bound by any particular theory of operation, it is believed
that LHC in
the light fraction 92 includes compounds generated through upgrading reactions
in the primary
reactor 24, etc., in addition to LE-IC compounds that may be native to the HCO
and any that may
have been added in the premix 50, 50' and/or reactor 24, 24', 24" sections to
promote
disaggregation reactions. Additionally, it is believed that the HHC1 94
flowing from FDC1 90
includes compounds that either have undergone varying degrees of upgrading
upstream from
I-DC1 90 but which on the whole may be susceptible to further upgrading. As
such, they are the
appropriate object of further upgrading in the supplemental reactor section
80. Due to their
higher molecular weight (MW) compared with LHCs, the viscosity of HHC1 94 will
vary
directly as a function of LHCs removed by FDC1 90. Therefore, the operation of
FDC1 90
should be optimized to ensure facile flow of HHC1 94 in a downstream manner
from FDC1 90
and prevent HHCs from becoming so concentrated and highly associated as to
reduce their
proclivity to undergo further upgrading reactions in the supplemental reactor
section 80.
Turning again to Figure 12, HHC1 94 flowing downstream through premixer 50' to
the
supplemental reactor section 80 is heated progressively in a particular
embodiment to a
predetermined maximum temperature T(max)2 before reaching the outlet of the
same, which
temperature is between about 1.0 and 1.1 times T(max)1, the predetermined
maximum
temperature to which the reaction mixture is heated in the primary reactor
section. Though not
wishing to be bound by a particular theory of operation, in this embodiment
the principal
objective is to further drive Type I reactions, described previously in this
specification, to an
extent beyond that which might have been achieved in the primary reactor
section, yet without
fostering appreciable upgrading by means of Type 11 reactions in addition to
Type I reactions.
Given that much or most of the water present in the product mixture flowing
into FDC1 90 is
removed in the light fraction 92, HHC1 94 may be contacted with water or steam
(such as
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recycled from light fraction 92) at one point or a plurality of points
downstream of FDC1 94, e.g.
in the supplemental reactor section 80 and/or the optional premix section 50'
as shown.
In another particular embodiment, the supplemental reactor section
progressively raises
the reaction mixture to a predetermined maximum temperature T(max)2 of between
about 1.1
and 1.4 times T(max)1 where Type II reactions are believed to occur. Given
that such reactions
are thought to be thermolytic instead of involving Type I hydrolytic
mechanisms, the contacting
of the HHC1 by water and/or steam in the premix section 50' and/or the
supplemental reactor
section 80 is optional. Nevertheless, said contacting may offer benefits aside
from the possibility
to promote further upgrading by Type I mechanisms. First, given that the
thermal conductivity of
.. water may be about 3-4 times higher than that of hydrocarbons including
bitumen, the presence
of just 5% to 10% water in the reaction mixture may increase the efficiency of
thermal energy
transfer from the walls of the shell-and-tube reactor to the HHC residue
flowing therethrough,
which is believed to minimize or prevent coke formation that would otherwise
occur, for
example at the walls of the shell-and-tube reactor. Additionally, not being
bound by a particular
theory of operation, the addition of water in controlled amounts may serve
helpfully to prevent
HHC components from becoming so concentrated and highly associated as to
reduce their
proclivity to undergo further upgrading reactions in the supplemental reactor
section 80.
However, as has been explained in detail elsewhere, the heat capacity of water
is approximately
twice that of hydrocarbons. Thus, in a particular embodiment of the present
invention, the
.. economics of HHC residue upgrading in the supplemental reactor section may
be optimized by
taking into consideration the quality and quantity of HHC residue fed
thereinto; the quality and
quantity of the upgraded hydrocarbon product flowing therefrom; and the amount
of thermal
energy applied, where the latter depends on 1) controlling the amount of LHC
removed in FDC1
90 and the amount of water or steam mixed with the HHC residue flowing through
the
.. supplemental reactor section 80; and 2) the predetermined temperature
T(max)2 to which the
reaction mixture is heated.
Compared with the approach described earlier in this specification, at least
four benefits
may accrue through the removal of LHC from the reaction mixture flowing into
FDC1 90 from
the primary reactor section 24, 24', 24". First, it results in a reduction of
the quantity of material
that must be heated to a temperature of T(max)2. In the particular case where
T(max)2 equals
about 1.1 to 1.4 times T(max)1, this reduction results in a corresponding
reduction in thermal
energy that must be invested to heat the HHC residue 94 compared with the case
where nol-DC1
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90 is used to reduce LHC. Second, corresponding to this reduction in material
mass is a
reduction in material volume, which reduces the hydraulic capacity, or
throughput, required of
the supplemental reactor section 80. This reduces not only the energy input
required, but the size
and cost of the capital equipment associated with the supplemental reactor
section. Third,
regardless of the value of T(max)2, removal by FDC1 90 of LHC in the product
mixture 34
avoids subjecting components in the same to higher-severity conditions, which
may cause them
to undergo further reactions to yield undesirable, lower-molecular-weight
products. Finally, by
way of corollary with the preceding benefit, removal of LHC from the product
mixture 34
provides the possibility to improve selectivity of the upgrading process
toward those components
that do not initially yield readily to upgrading under lower-severity
conditions.
In a particular embodiment, the product mixture 34' exiting the outlet of the
supplemental
reactor section 80 may flow into a recovery section 38, 52, etc., such as
described hereinabove
with respect to Figs. 1-6B and 11. Turning once again to Figure 12, in another
particular
embodiment the product mixture flows optionally into a second flash
distillation column 96, also
designated as FDC2. As with FDC1, FDC2 obtains two fractions from the product
mixture
flowing thereinto from the supplemental reactor section: a light fraction 98
including water and
light hydrocarbons, the latter being designated LHC2; and a heavy fraction
100, designated
HHC2. As in the case of FDC1, these two fractions are differentiated from each
other on the
basis of the boiling point of the light fraction, in this case being
designated BP(LHC2), which in
one embodiment is below about 280 C, while in yet another particular
embodiment BP(LHC2)
may be below about 220 C, and in yet another particular embodiment, it may be
less than about
160 C. LHC1 and LHC2 may be combined to obtain a single LHC product stream
(not shown).
Referring again to Figure 12, in a particular embodiment a portion of HHC2 100
may be
optionally recycled, e.g., at 102, to an upstream point in the supplemental
reactor section, e.g. to
the premix section 50', where it is combined with HHC1 94. And in yet another
embodiment, a
portion of the light hydrocarbons from 92 and/or 98, may in particular
embodiments be
optionally recycled and mixed with HCO feed 22 at points in the primary
reactor section 24, 24',
24" before the reaction mixture reaches a temperature of about 60% to 70% of
T(max)1, and/or
in the premix section 50. Though thought to be desirable to minimize the
amount of LHC that is
subjected to higher-severity upgrading conditions, the mixing of limited
amounts of LHC with
HHC1 in the premix section 50 may beneficially permit reduction in the amount
of water
required to ensure that heavy components in the reaction mixture flowing
through the
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supplemental reactor section 80 are sufficiently disaggregated, the benefit
including a
corresponding reduction in the energy required to heat the reaction mixture to
T(max)2. Though
not wishing to be bound by a particular theory of operation, it is believed
that components in
HHC2 100 may undergo further upgrading if again subjected to the upgrading
conditions in the
supplemental reactor section 80 by recycling at 102. As in the case of HHC1
94, this approach
enables the additional, selective upgrading of heavy components while
minimizing or avoiding
the subjecting of fighter components 98, removed in FDC2 96, to higher-
severity conditions in
the supplemental reactor section. This helps to minimize the risk of
undesirable decomposition of
LHCs that might reduce overall upgrading yield and provides the means to
promote further
upgrading without increasing the predetermined maximum temperature T(max)2 in
the
supplemental reactor section.
Referring now to Figure 13, in another particular embodiment, product mixture
outputs
from the primary reactor subsection 24, 24', 24' 'and supplemental reactor
subsection 80 feed
into a single flash distillation column 96, while a portion of the HHC residue
100' is recycled at
102 to serves as a feed to the supplemental reactor section 80, e.g., via
premix 50' as shown. By
way of example, suppose that the flash distillation column 96 is operated in
such a way that the
HHC residue includes components whose boiling points are greater than about
280 C, and that
the LHC 98' and HHC residue 100' each represents about 50% of the combined
product mixture
flowing into the flash distillation column 96. In a particular embodiment the
supplemental
reactor has the capacity to process up to about 80% of the HHC residue, and
therefore would
have a throughput of about 40% that of the primary reactor section. Again, the
benefit of this
processing scheme is that it provides the possibility to optimize the
economics of upgrading
through the control of multiple process variables, including but not limited
to the maximum
predetermined temperatures T(max)1 and T(max)2; the operation of the flash
distillation column
to determine the maximum boiling point of components in the LHC fraction,
BP(LHC); and the
residue recycle rate (the fraction of the residue that is processed through
the supplemental reactor
section). Other variables, discussed elsewhere, include flow rate, the amounts
of LHC and water
or steam in the reaction mixture flowing through the primary reactor section
(or subsections);
and the size and number of thermal soak chambers, which in combination with
flow rate
determine the length of time that reaction mixtures in the primary or
supplemental reactor
sections experience the predetermined maximum temperatures T(max)1 and
T(max)2,
respectively.
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Turning now to Figure 14, the reaction mixture proceeds in a linear sequence
from the
premix section 50 and then through the primary reactor section 24, 24', 24",
the optional
supplemental reactor section 80, including optional thermal soak chambers 74.
Finally, the
product mixture flows into a flash distillation column 96, which, as in the
cases of FDC1 and
FDC2 described previously, yields LHC 98' and HHC residue 100' products
streams whose
components are differentiated on the basis of boiling point. In a particular
embodiment a portion
of the HHC residue 100' from the flash distillation column is recycled at 102'
to the premix
section 50 disposed upstream from the primary reactor section where it
combines with HCO feed
22 and with LHC (e.g., recycled from 98') and/or water. In another particular
embodiment
depicted in Figure 15, the residue recycle 102 instead flows into an optional
premix section 51'
disposed between the primary reactor section 24, 24', 24" and supplemental
reactor section 80
where it combines with the reaction mixture flowing from the primary reactor
section, and
optionally with LHC and/or water supplied to premix section 50. In contrast
with embodiments
depicted in Figures 12 and 13, the capacity of the supplemental reactor
section 80 corresponding
to the embodiment shown in Figure 14 is substantially the same as that of the
primary reactor
section 24, 24', 24", whereas the capacity of the supplemental reactor section
used in Figure 15
would be larger in order to handle the recycled residue 100'.
As discussed above, in the embodiments depicted in Figures 12 ¨ 15, flash
distillation
columns yield two fractions: a light fraction including water and light
hydrocarbons (LHCs), and
a HHC fraction referred to alternatively as HHC residue or simply as residue,
which includes
hydrocarbons that are differentiated from the LHCs on the basis of boiling
point: LHC
components have boiling points below about BP(LHC) while HHC components have
boiling
points that are higher than about BP(LHC). Though the flash distillation
columns serve to
separate LHCs from HHC residue on the basis of boiling point, that separation
is not absolute.
Rather, some compounds with boiling points below BP(LHC) remain in the HHC
residue and
some with boiling points above BP(LHC) distill with the LHC. In consideration
of process
optimization objectives, the value of BP(LHC) and the extent of overlap
between the two curves
(Fig. 16) are determined by (i) the design of the distillation column,
including the number of
theoretical plates; (ii) the flow rate of the product mixture into the column;
and (iii) the
temperature of the product mixture flowing into the column; and (iv) the
operating pressure
inside of the column. Referring to Figure 16, BP(LHC) effectively represents a
temperature
range in which compounds contained in both LHC and H_HC residue distill, the
nominal value of
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BP(LHC) corresponding to the intersection 110 of the distillation yield curves
for LHC and HHC
residue, shown at 112 and 114, respectively. The process optimization
objectives may consider
among other things the yield and quality of LHC desired and the composition of
the HHC
residue. Not wishing to be bound by a particular theory of operation, the
flash distillation
columns in the embodiments depicted in Figures 12-15 should be controlled so
that the quantity
of LHC removed is not so great as to either permit the re-aggregation of
higher MW compounds
in the HHC residue to such an extent that they would resist subsequent
upgrading, or cause the
viscosity of the HHC residue to become intractably high.
A person possessing common knowledge regarding the design and operation of
distillation columns will recognize that flash distillation columns provide a
high-throughput,
coarse separation between components in a mixture on the basis of their
boiling points, in
contrast with distillation columns that provide sharper resolution in regard
to boiling point by
dint of having more theoretical plates. However, the repeated reference in
this specification to
the use of flash distillation columns is not intended to limit the optional
use of higher-resolution
columns to serve a process optimization goal such as minimizing the range of
the boiling point
overlap between heavy and light fractions. Similarly, the portrayal in Figures
12-15 of flash
distillation columns with only two cutpoints for obtaining light and heavy
fractions, e.g.
water/LHC and HHC residue, respectively, is not meant to preclude the optional
configuration of
a distillation column to have three or more cutpoints to obtain from a product
mixture two or
more LHC fractions in addition to the HHC residue. An additional feature of
the process whose
importance would be understood by those skilled in the art, and therefore not
detailed in the
present specification, is the removal from LHC of any relatively high-
volatility compounds that
might accompany LHC and water in the light fraction obtained by the flash
distillation column,
which removal may be desirable to make the LHC suitable for purposes described
herein.
As already has been discussed in some detail with respect to embodiments
represented in
Figures 12 and 13, the embodiments depicted in Figures 12-15 afford
possibilities to optimize
the upgrading process. Referring now to Figure 17, in a particular embodiment
of the present
invention all LHC and HHC residue outputs are combined to produce a type of
synthetic crude
whose API, distillation yield, and other properties resemble those of
conventional crude and
make it pipelineable and suitable as a refinery feedstock. The objective of
optimization is to
maximize efficiency, a circumscribed view of which, by this embodiment,
considers only the
quantity and value of process outputs obtained from a given quantity and value
of inputs that
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principally include energy and HCO. In a particular embodiment, the upgrading
process is
designed and operated at a severity such that substantially all LHC and HHC
residue products
obtained from HCO feed of a given quality are consumed to produce a synthetic
crude oil whose
properties meet but do not exceed the minimum values that define
pipelineability. Embedded in
this definition of efficiency is the concept of yield, which is understood in
terms of the amount of
high-value, pipelineable material produced relative to HCO and thermal energy
input into the
process. This efficiency objective may be achieved by a number of means
including but not
limited to controlling (i) values for the predetermined maximum temperatures
T(max)1 and
T(max)2 to minimize the net energy applied per-barrel for upgrading and/or
(ii) the residue
recycle ratio (the fraction of HHC residue exiting the bottom of a flash
distillation column that is
directed to a point upstream from the flash distillation column where it is
combined with the
reaction mixture) and/or (iii) the number and size of thermal soak chambers
interposed in the
flow path.
The constrained view of efficiency just described may be useful to a point,
but it is also
incomplete. Another particular embodiment takes a holistic view of efficiency
that includes all
aspects of the constrained view, but also gives consideration to additional
negative-value inputs
and outputs, examples of which include coke (also referred to as petroleum
coke, pet coke, or
petcoke when derived from petroleum) and light gases such as methane, ethane,
and propane, all
of which are generated in significant quantities by the HCO upgrading
processes discussed
hereinabove. These by-products have value as fuel that may be used to generate
thermal energy
required to drive the upgrading process. However, greenhouse gas emissions per
unit energy
produced by petcoke is similar to that for coal. The additional production of
pollutants such as
SO2 and NOx further exacerbates the problem. Thus, production of coke both
reduces the yield
of positive-value upgraded HCO and disproportionately increases the amount of
negative-value
emissions.
Referring now to Figures 18 and 19, in other particular embodiments the
unconstrained
definition of HCO upgrading efficiency expands to include production of LHC in
excess of that
required to produce pipelineable synthetic crude oil as described above in
relation to Figure 17.
In these particular embodiments of the present invention, processes depicted
in Figures 12-15 are
designed and operated so as produce said excess quantities of LHC by, for
example, designing
and controlling the upgrading process to operate at higher severity while also
minimizing coke
formation, the designing and controlling being accomplished through means
including but not
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limited to (i) increasing values for the predetermined maximum temperatures
T(max)1 and
T(max)2 in the primary and supplemental reactor sections, respectively; and/or
(ii) interposing a
plurality of thermal soak chambers in the flow path upstream from flash
distillation columns;
and/or (iii) increasing the size of said chambers to increase the residence
time that the reaction
mixture experiences upgrading conditions; and/or (iv) increasing the residue
recycle ratio.
Excess LHC also may be generated by designing and operating flash distillation
columns so as to
increase the amount of LHC recovered from the reaction mixture while
respecting previously
discussed cautions against removing too much LHC in flash distillation
columns.
Referring again to Figure 18, excess LHC obtained by such means is, in a
particular
embodiment, blended with HCO that has not been upgraded to produce diluted
bitumen (dilbit),
which is another type of upgraded HCO that is pipelineable. This dilbit
production can be
performed by means of blending operations that are either proximate to or
distant from the
location of the hydrothermal upgrader, the LHC being transported in the latter
case by truck, rail,
or pipeline.
Turning again to Figure 19, all or a portion of the excess LHC may be blended
with any
combination of (i) HHC residue generated by an upgrading process such as those
depicted in
Figures 12-15, and (ii) HCO that has not been subjected to upgrading. In a
particular
embodiment, substantially all of the excess LHC is mixed with all HHC residue
generated by the
upgrading process and with an amount of HCO corresponding to the maximum that
can be
.. combined with the other blending components to produce synthetic crude oil
that meets the
requirements for pipelineability. In yet another embodiment, excess LHC not
consumed in the
production of synthetic crude oil may be used for dilbit production, whether
in proximity to or at
a location remote from the point of LHC production.
The invention described heretofore with reference to specific exemplary
embodiments for
the purposes of illustration and description is not intended to be exhaustive
or to limit the
invention to the precise form disclosed. Many modifications and variations are
possible in light
of this disclosure. For example embodiments of the present invention
corresponding to
depictions in Figures 12-15 may be provided with a recovery section wherein
some portion of
the thermal energy invested into the process may be captured from the product
mixture stream or
.. from the LHC/water and HHC residue product streams and used to facilitate
the heating of the
reaction mixture upstream from the flash distillation column, such as shown
and described
hereinabove with respect to Figs. 1-6B and 11. Similarly, the embodiments of
Figures 12-15 may
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supply water, e.g., which is separated from LHC after removal from the flash
distillation column,
to the HHC residue product stream 94, 100, to remove undesirable water-soluble
materials
therefrom, which may include compounds of vanadium and nickel. Also, it should
be recognized
that water removed from the LHC downstream of the flash distillation column(s)
may be
recovered and fed back to the HCO or the reaction mixture at an upstream point
in the flowpath
of the reaction mixture, such as shown and described with respect to Figs. 2-
6B. Moreover, it
should be understood that a controller 32 (Fig. 1) may be used to control
operation, including
optimizing the blending of HHC residue and LHC, or of HHC residue, LHC, and
HCO, or of
LHC and HCO to produce synthetic crude oil or dilbit that meets the minimum
requirements of
pipelineability as described herein. It is intended that the scope of the
invention be limited not by
this detailed description, but rather by the claims appended hereto.
The understanding of process efficiency now has been expanded to encompass
upgraded
HCO obtained by means of embodiments depicted in Figures 12-15, through the
simple dilution
with excess LHC as depicted in Figures 18 and 19, and by blending with HHC
residue and
excess LHC according to Figure 19. Consider now that the SAGD (Steam Assisted
Gravity
Drainage) oil recovery process used widely in Alberta is known to be energy
intensive, owing
primarily to the generation and downhole injection into oil sands deposits of
large quantities of
high pressure steam to produce water-bitumen emulsions that are recovered by
pumping to the
surface where water is separated to yield HCO. Extraction by the SAGD process
of bitumen in
oil sand deposits involves complex thermodynamic, physical chemistry, and mass
transfer
phenomena which others have shown can be advantageously altered by addition
into the steam of
hydrocarbon, the core principle being simply that the bitumen in the oil sands
does not mix
efficiently with water but does with added hydrocarbons. The enhancement is
the increasing of
the yield of recovered HCO per unit energy invested, one measure being the
cumulative steam-
to-oil ratio (CSOR), which is the amount of steam as liquid water that must be
generated and
injected downhole to produce a barrel of HCO. Currently in the range of 2:1 to
3:1, a solvent-
enhanced SAGD (SAGD/SE) process has the potential to reduce the CSOR by 25% to
50%,
resulting in a corresponding reduction in energy consumption and greenhouse
gas emissions
associated steam generation.
Generally, hydrocarbons employed to recover HCO by means of SAGD/SE are
"imported" to the wellhead. That is, they are not generated at the wellhead,
but instead must be
purchased in the marketplace and transported to the wellhead. Because cost is
an issue, and
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lower molecular-weight hydrocarbons such as propane or butane generally are
cheaper than
equivalent quantities of pentanes, gasoline, or distillates, economics of
SAGD/SE dictate use of
the former. A stated advantage of doing so is that HCO produced by SAGD/SE has
API values
that typically are in the range of 13-17, compared with values that are
typically below 10 for the
conventional SAGD process. However appealing those numbers may be, they belie
the reality
that conventional SAGD/SE is selective against heavier hydrocarbons that also
have value. In
effect, SAGD/SE that uses light, three- and four-carbon solvents actually may
have a lower yield
than conventional SAGD, yield being defined in terms of the amount of HCO
removed from oil
sand deposits relative to the total available HCO that is resident in those
deposits. Though not
wishing to be bound by a particular theory concerning mechanisms by which
SAGD/SE
operates, the higher API values for the HCO obtained by it are consistent
with the fact that
compounds in crude oils whose polarity, MW, and aromaticity are high, e.g.
asphaltenes, have
relatively low solubility in low-MW alkanes such as n-pentane, much less
propane and butanes,
compared with hydrocarbons that have both higher MW and diverse chemical
functionality.
Accordingly, SAGD/SE employing low-MW alkanes may be selective against
compounds that
otherwise have fuel value and may readily form emulsions with water to enable
their recovery.
Of course, the relatively lower level of heavy components in the HCO recovered
by SAGD/SE is
an advantage when no convenient, economical means exists to upgrade those
components, such
as that afforded by the present invention. Thus, the purported advantage of
conventional
SAGD/SE comes at the expense of reduced overall recovery from oil sands
deposits of
hydrocarbons contained therein.
In a particular embodiment of the present invention, excess LHC produced as
described
hereinabove is injected with steam into oil sand deposits. Though not wishing
to be bound by a
particular theory of operation, such an LHC-enhanced SAGD process (SAGD/LHC)
is thought
to have an important advantage over SAGD/SE that relies on propane, butanes,
or pentanes. The
chemical composition LHC is thought to include compounds whose molecular
weights and
molecular weight ranges are higher, and which have diverse chemical structures
including but
not limited hydrocarbons containing various straight-chain, branched, and
aromatic moieties. As
cess LHC generated by means of the ins afford higher yields than SAGD/SE that
relies on
aliphatic hydrocarbons with just 3-4 or even 5 carbons. Of some concern is the
possibility for
fugitive emissions of light hydrocarbons such as propane from oil sand
deposits where
SAGD/SE is being applied, such emissions being considered greenhouse gases. By
contrast, such
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emissions from SAGD/LHC process are expected to be substantially lower due to
the lower
vapor pressure of LHCs.
Also, the recovery of low MW solvents used for SAGD/SE, which are contained in
the
water-HCO mixture extracted from oil sand deposits, is an economic necessity,
as doing so
reduces the amount of solvent that must be purchased and transported to the
wellhead.
Additionally, pipelineability specifications relating to vapor pressure
effectively limit the
quantities of propane and butanes that may be contained in crude oils. Yet,
the recovery and
reuse of such solvents relies on equipment for their handling, purification,
and storage,
representing an additional capital burden. By contrast, LHC used in SAGD/LHC
and recovered
with HCO is compatible with pipelineable crude oil and therefore does not need
to be removed,
e.g. when dilbit is produced from HCO. Equally important, LHC used in SAGD/LHC
does not
need to be purchased on the open market and transported to the wellhead, but
instead can be
produced by means of the instant invention when installed and operating in
relatively close
proximity to the wellhead.
In yet another particular embodiment, excess LHC generated by means of the
instant
invention may be transported by truck, rail, or pipeline to a site where
SAGD/SE is being used to
extract HCO from oil sand deposits, the purpose being to supplement the low-MW
hydrocarbons
conventionally used in SAGD/SE, or to replace the same in the event that
supplies thereof are
limited, or if market value of the same, relative to that of LHC, fluctuates
in a manner that
warrants such a hybridized strategy. Thus, this particular embodiment serves
to mitigate market-
based risks associated with a one-dimensional SAGD/SE approach.
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In the preceding specification, the invention has been described with
reference to specific
exemplary embodiments for the purposes of illustration and description. It is
not intended to be
exhaustive or to limit the invention to the precise form disclosed. Many
modifications and
variations are possible in light of this disclosure. It is intended that the
scope of the invention be
limited not by this detailed description, but rather by the claims appended
hereto.
It should be further understood that any of the features described with
respect to one of
the embodiments described herein may be similarly applied to any of the other
embodiments
described herein without departing from the scope of the present invention.
Having thus described the invention, what is claimed is:
34
