United States Patent [19]
Bain et at
[11] Patent Number:
[45] Date of Patent:
4,778,586
Oct. 18, 1988
[57] ABSTRACT
28 Claims, 1 Drawing Sheet
1189011 6/1985 Canada.
420650 3/1974 U.S.S.R..
OTHER PUBLICAnONS
"Conversion of Petroleum", A. N. Sachanen, D. Sc.,
Reinhold Publishing Corporation, New York, 1948.
"Visbreaking ... Still the Basic Process for Fuel Reduction",
M. G. Boone and D. F. Ferguson, The Oil and
Gas Journal, Mar. 22, 1954.
"Visbreaking: A Flexible Process", A. Rhoe and C. de
Blignieres, Hydrocarbon Processing, Jan. 1979.
"Visbreaking: More Feed for FCC", R. Hournac, J.
Kunn and M. Notarbarto10, Hydrocarbon Processing,
Dec. 1979.
"Visbreaking Process has Strong Revival", Martin Hus,
Technology Oil and Gas Journal, Apr. 13, 1981, vol. 79.
"Rebuilding Hydrocarbons", W. L. Nelson, Petroleum
Refinery Engineering, 4th Ed., McGraw-Hill Book
Company, Inc., 1958.
Primary Examiner-Paul E. Konopka
Attorney, Agent, or Firm-Sheridan, Ross & McIntosh
A method is disclosed for improving the transportability
of a hydrocarbon composition by passing an influent
feed stream of composition into a downcomer to provide
a hydrostatic column of fluid. The influent stream
is heated by heat exchange with an effluent product
stream wherein at least one of the streams is in turbulent
flow. The feed stream is pressurized by the hydrostatic
pressure ,head to a reaction pressure of at least about
1000 psi.I The heated and pressurized feed stream is
contacted with an active heat source in a reaction zone
to increase the temperature of the feed stream to a reaction
temperature of between about 300· C. and the coking
temperature of the hydrocarbon composition. The
temperature differential between the active heat source
and the feed stream in the reaction zone is maintained at
less than about 30· C. to provide a treated effluent
stream which is brought into heat exchange contact
with the influent stream. The treated composition has a
lower viscosity than the feed composition.
Related U.S. Application Data
Continuation-in-part of Ser. No. 771,205, Aug. 30,
1985, abandoned.
Int. Cl.4 ClOG 9/14; ClOG 9/00
U.S. Cl 208/132; 137/13;
165/45; 196/110; 208/106; 208/125
Field of Search 208/106, 125, 131, 132;
196/110; 137/13; 165/45
[56] References Cited
U.S. PATENT DOCUMENTS
1,479,653 1/1924 Davidson 196/65
1,828,691 10/1931 Stransky et al. 196/65
2,135,332 11/1938 Gary 196/62
2,160,814 6/1939 Arveson 196/50
2,293,421 8/1942 Baetz 196/110
2,587,703 3/1952 Deansely 196/65
2,651,601 9/1953 Taff et al. 196/73
2,695,264 11/1954 Taff et al. 196/50
2,752,407 6/1956 Cahn 266/683
2,818,419 12/1957 McKinley et al. 260/451
2,844,452 7/1958 Hasche 48/196
2,862,870 12/1958 Voorhies 208/56
2,900,327 8/1959 Beuther 208/106
2,937,987 5/1960 Jenkins 208/108
2,981,747 4/1961 Lang et al. 260/451
3,156,642 11/1964 Trantham et al. 208/120
3,170,863 2/1965 Spillane et al. 208/3
(List continued on next page.)
FOREIGN PATENT DOCUMENTS
1184523 3/1985 Canada.
[51]
[52]
[58]
[63]
[54] VISCOSITY REDUctION PROCESSING AT
ELEVATED PRESSURE
[75] Inventors: Richard L. Bain, Golden; John R.
Larson, Boulder; Dennis D.
Gertenbach, Golden; Daniel W.
Gillespie, Wheatridge; Joseph J.
Leto, Broomfield, all of Colo.
[73] Assignee: Resource Technology Associates,
Boulder, Colo.
[21] .Appl. No.: 58,881
[22] Filed: Jun. 5, 1987
4,778,586
Page 2
U.S. PATENT DOCUMENTS
3,306,839 2/1967 Vaell 208/59
3,310,109 3/1967 Marx et al. 166/7
3,320,154 5/1967 Tokuhisa et al. 208/130
3,412,011 11/1968 Lindsay 208/113
3,439,741 4/1969 Parker 166/372
3,442,333 5/1969 Meldau 166/272
3,523,071 8/1970 Knapp et al. 208/14
3,738,931 6/1973 Frankovich et al. 208/67
3,767,564 10/1973 Youngblood et al. 208/92
3,775,296 11/1973 Chervenak et al. 208/108
3,803,259 4/1974 Porchey et al. 208/106
3,948,755 4/1976 McCollum et aI 208/113
3,989,618 11/1976 McCollum et al. 208/208 R
4,042,487 8/1977 Seguchi et at. 208/48 R
4,089,340 5/1978 Smith et al. 137/13
4,248,306 2/1981 VanHuisen et al. 166/305 R
4,252,634 2/1981 Knulbe et al. 208/48 R
4,298,457 11/1981 Oblad et al. 208/107
4,334,976 6/1982 Van 208/8
4,354,922 10/1982 Derbyshire et al. 208/68
4,379,747 4/1983 Van 208/251 H
4,428,828 1/1984 Bose 208/208 R
4,432,864 2/1984 Myers et at. 208/120
4,448,665 5/1984 Zaczepinski et al. 208/8
4,460,012 5/1984 Murthy et al. 208/130
4,465,584 8/1984 Effron et at. 208/56
4,469,587 9/1984 Tailleur et al. 208/61
4,478,705 10/1984 Ganguli 208/59
4,481,101 11/1984 Van 208/50
4,508,614 4/1985 Van 208178
4,560,467 12/1985 Stapp 208/89
4,631,384 12/1986 Cornu 291/121
4,671,351 6/1987 Rappe 165/45
4,741,386 5/1988 Rappe 165/45
u.s. Patent Oct. 18, 1988 4,778,586
1
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2
phalting to mild visbreaking to severe thermal cracking.
Distillation and deasphalting processes result in separation
of the heavy portion of the oil, i.e. the residuum,
from the remaining lighter portion, with only the ligh-
5 ter end being transported.
A number of processes which involve heating a
heavy oil to improve its transportability have been tried
over the years. A thermal treating process to reduce the
viscosity and improve transportation of the oil has been
10 disclosed by Engle in U.S. Pat. No. 3,496,097 (1970).
This process involves heating the oil between 500° F.
and 700° F. for at least 24 hours. The process has the
disadvantage of being time and energy consumptive and
producing substantial amounts of gas which are not
15 readily used in the field.
Scott et al. in U.S. Pat. No. 3,474,596 (1969) describe
a process for reducing the viscosity of a stream of viscous
fluid flowing within a pipeline by diverting a portion
of the stream and heating it to about 850° F. to 900°
F. (454° C.-482° C.) and 200 to 400 psig at which thermal
degradation or "visbreaking" of at least some of the
constituents thereof takes place. This heated portion is
then blended with the remainder of the stream to reduce
the viscosity of the bulk material. This process, however,
only modifies a portion of the oil. Additionally,
that portion which is modified must be taken from the
fraction of "dry oil" which is obtained from a crude
oil-water separator.
Huang in U.S. Pat. No. 4,298,455 (1981) discloses that
the pumpability of a heavy hydrocarbon oil, such as a
crude, reduced crude or other oil with an API gravity
ofless than 15" is improved by using a viscosity reducing
or visbreaking heat treatment. The disclosed process
involves heating the oil at between 800° F. and 950°
F. (427° C.-51O° C.) between two and thirty minutes
and at a pressure of 100 to 1500 psig. To minimize the
amount of coke or tar and gas formed during this visbreaking
process, the visbreaking is carried out in the
presence of a chain transfer agent and a free radical
initiator. This process requires the careful control of the
concentration of the initiator and transfer agent in conjunction
with adjustment of the residence time at reaction
temperature to minimize coke formation.
A method which involves reducing the viscosity and
sulfur content of a heavy crude as it is being produced
is disclosed by Meldau in U.S. Pat. No. 3,442,333 (1969).
This method involves injecting steam at the wellhead
through a conduit which extends down-hole. The steam
50 heats the oil to a temperature in the range of 550°
F.-700°F. (288° C.-371° C.). The rate of production of
the oil is controlled so that the oil is at temperature
within the well for at least 24 hours. This process has
the disadvantages of long contact times at temperature,
high energy requirement, low production rates, and the
necessity for special equipment in each well-hole.
A form of thermal cracking known as visbreaking is
well known in the art. As disclosed by Biceroglu et al.
in U.S. Pat. No. 4,462,895 (1984), visbreaking conditions
can include temperatures from 750° F.-950° F.
(399° C.-51O° C.) and pressures of 50-1500 psig. Other
conditions disclosed include a temperature of 850°
F.-975° F. (454° C.-524° C.) and a pressure of 50-600
psig. Beuther et al. U.S. Pat. No. 3,132,088 (1964). Normally
the residue from "topped" or "reduced" crudes is
the feedstock for refmery visbreaking operations. Taff
et al. U.S. Pat. No. 2,695,264 (1954). It has been disclosed
by Beuther et al. in U.S. Pat. No. 3,324,028
4,778,586
1
VISCOSITY REDUCTION PROCESSING AT
ELEVATED PRESSURE
BACKGROUND OF THE INVENTION
Development of many of the world's petroleum re- 20
serves is hindered or prevented by the nature of crude
oil where the viscosity, pour point and API gravity
renders the crude oil unsuitable for pipeline transportation.
Varied methods of producing pipeline-quality oil
from such crudes have been used. In general, such 25
methods can be categorized as either physical or chemical
treatments.
Physical treatments change the physical properties of
the oil to produce a pumpable fluid, but do not change
the chemical composition of the oil itself. As discussed 30
by Flournoy et al. in U.S. Pat. No. 4,134,415 (1979) a
common method involves dilution of the heavy crude
with lighter fractions ofhydrocarbons. This can involve
the use of large amounts of expensive solvents to transport
a relatively cheap product and requires the avail- 35
ability of the diluent which can be inconvenient in certain
oil fields. Another method disclosed by Flournoy et
al. involves he!lting the heavy oil to reduce its viscosity.
This method requires the installation of heating equipment
along the pipeline and insulation of the pipeline 40
itself. Such a procedure is expensive and uses a large
amount of energy. The extent of decrease in viscosity
which can be achieved by an increase in temperature
varies widely between heavy oils depending on the oil
composition. Such physical treatments do not upgrade, 45
i.e. enhance the value of, the oil and, in fact, usually
increase the overall cost ofoil processing. Nevertheless,
physical treatments provide a simple solution and are
most widely used today. In many applications, dilution
with lighter crudes is coupled with pipeline heating for
pumping very heavy crudes. It is also possible to add
water to reduce the pressure gradients as discussed by
B. L. Moreau in an article "The Pipeline Transportation
of Heavy Oils", The Journal of Canadian Petroleum
Technology, p. 252, 1965. However it is difficult to main- 55
tain proper flow in this system and still obtain the desired
viscosity reduction. Other methods such as the
addition of surfactants to form oil-in-water emulsions
have been used. Flournoy et aI., U.S. Pat. No. 3,943,954
(1976). 60
Chemical treatments can involve contacting the oil
with a strong base to form an oil-in-water emulsion
which is more easily transported. Kessick et aI., Canadian
Pat. No. 1,137,005 (1982). However, chemical
treatments typically require changing the hydrogen to 65
carbon ratio of the oil, either by reducing the carbon
content or by addition of hydrogen. Carbon reduction
technologies range from simple distillation and deas-
CROSS-REFERENCE TO RELATED
APPLICATION
This application is a continuation-in-part of copending
and commonly assigned U.S. patent application Ser.
No. 771,205 filed Aug. 30, 1985 now abandoned.
FIELD OF INVENTION
This invention relates to a method for improving the
transportability of heavy oils and other hydrocarbons
by reducing viscosity in order to render them more
suitable for transportation by pipeline and ship and/or
to provide enhanced value for refinery processing to
increasing the API gravity.
4,778,586
4
before passing it to a visbreaking heater. Black (supra)
teaches that it is desirable to minimize vaporization
during cracking to maintain only a liquid phase. Black
used mechanical pressure of up to 1000 psi and the
addition of a liquid diluent to maintain the liquid phase.
In view of the disadvantages of the processes described
hereinabove, there is a need for a process suitable
for well-site locations by which viscous crudes can
be rendered more pumpable. More particularly, it
10 would be advantageous to have a process which, unlike
traditional visbreaking, is suitable for untopped, rather
than topped, feeds and which uses lower temperatures
to achieve the same or greater viscosity reductions.
It has now been found that significant reductions in
the viscosity of heavy hydrocarbon mixtures can be
attained with a process using a vertical tube reactor.
Vertical tube reactors which oridinarily involve the use
of a subterranean U-tube configuration for establishing
a hydrostatic column of fluid sufficient to provide a
selected pressure are known. This configuration provides
a less expensive way to achieve high pressures
than with standard high pressure pumps. This type of
reactor has been primarily used for the direct wet oxidation
of materials in a waste stream and particularly for
the direct wet oxidation of sewage sludge.
Bower in U.S. Pat. No. 3,449,247 discloses a process
in which combustible materials are disposed of by wet
oxidation. A mixture of air, water and combustible material
is directed into a shaft and air is injected into the
mixture at the bottom of the hydrostatic column.
Lawless in U.S. Pat. No. 3,606,999 discloses a similar
process in which a water solution or suspension of combustible
solids is contacted with an oxygen-containing
gas. Excess heat is removed from the apparatus by either
diluting the feed with the product stream or withdrawing
vapor, such as steam, from the system.
Land, et al. in U.S. Pat. No. 3,464,885 (1969) is directed
to the use of a subterranean reactor for the digestion
of wood chips. The method involves flowing the
material through counter-current coaxial flow paths
within a well-bore while flowing heated fluid coaxially
of the material to be reacted. The reactants, such as
sodium hydroxide and sodium sulfate, are combined
with the wood chip stream prior to entry into the Utube
which is disposed within a well-bore.
Titmas in U.S. Pat. No. 3,853,759 (1974) discloses a
process in which sewage is thermally treated by limiting
combustion of the material by restricting the process to
the oxygen which is present in the sewage, i.e. no additional
oxygen is added. Therefore, it is necessary to
provide a continuous supply of heat energy to effect the
thermal reactions.
McGrew in U.S. Pat. No. 4,272,383 (1981) discloses
the use of a vertical tube reactor to contact two reactants
in a reaction zone. The method is primarily directed
to the wet oxidation of sewage sludge in which
substantially all of the organic material is oxidized.
There is heat exchange between the inflowing and
product streams. The temperature in the reaction zone
is controlled by adding heat or cooling as necessary to
maintain the selected temperature. It is disclosed that
when gas is used in the reaction, it is preferred to use a
series of enlarged bubbles known as "Taylor Bubbles".
These bubbles are formed in the influent stream and are
transported downward into the reaction zone. It is disclosed
that preferably air is introduced into the influent
stream at different points with the amount of air equaling
one volume of air per volume of liquid at each injec-
3
(1967) that resids and certain heavy crudes with an API
gravity below about 20· can be exposed to visbreaking
conditions. This patent, however, teaches that the resids
or crude should be hydrodesulfurized before visbreaking
at 800· F.-1000· F. (42r C.-538· C.) at pressures of 5
50-1000 psig. Such "visbreaking" processes are not
practical for in the field treatment of whole crude because
of the additional facilities required to pretreat the
feedstock and to recover and process products from the
treatment.
The principal variables in single-pass visbreaking
have been reported to be furnace outlet temperature,
residence time and pressure. Beuther et al., "Thermal
Visbreaking of Heavy Residues", The Oil and Gas Journal,
Vol. 57, No. 46, p. 151 (1959). An increase in any of 15
the three variables is said to result in an increase in
visbreaking severity. Shu et al. in U.S. Pat. No.
4,504,377 (1985) and Yan et al. in U.S. Pat. No.
4,522,703 (1985).
It has been disclosed that at higher severities there is 20
an increased tendency to form coke deposits in the
heating zone or furnace. Black in U.S. Pat. No.
1,720,070 (1929) teaches that operating at lower temperatures
for increased lengths of time provides "a much
smaller amount of carbon is deposited than is deposited 25
at higher temperatures." Hanna et al. in U.S. Pat. No.
1,449,227 (1923) disclose the continuous circulation of a
stream of oil from an evaporating chamber through a
heating coil to maintain the temperature of the oil in the
chamber at the desired cracking temperature. The tem- 30
perature differences between the oil in the chamber and
the heating coil is kept small to minimize cracking in the
coil. Hess in U.S. Pat. No. 1,610,523 (1926) teaches that
it is desirable to avoid local overheating in order to
prevent excessive coke formation in cracking systems of 35
oil distillation. Akbar et aI., "Visbreaking Uses Soaker
Drum", Hydrocarbon Processing, May 1981, p. 81 discloses
that, when there is a high temperature differential
between the tube wall in a furnace cracker and the bulk
temperature of the oil, the material in the boundary 40
layer adjacent to the tube wall gets overcracked. Therefore,
the coking rate is roughly a function of the inside
boundary layer temperature. In furnace cracking this
boundary layer is commonly 30· C.-40· C. higher than
the bulk temperature. In soaker cracking the skin tem- 45
perature is lower but still is reported to be above 480· C.
Therefore, the formation of coke is slower in a soaker
cracker but still causes regular shutdowns of the equipment
for coke removal.
Frequent shutdowns for coke removal from visbreak- 50
ing units can be tolerated in refinery operations where
there is adequate storage for the topped crude or residue
feedstock normally processed. However, this is
unacceptable in a field operation where crude is continually
produced and must be rapidly transported. Yan et 55
al. (supra) recognize the problem of coke formation.
They attempt to minimize the problem by adding "1-10
weight percent of finely divided solids in the heavy
hydrocarbon oil feedstream ... " in an attempt "... to
prevent the deposition of coke on the walls of the heat- 60
ing coils and reactor . . . "
Although some patents relating to visbreaking suggest
that whole crude can be used as a feedstock, this
has not proven possible with conventional processes
due to the pressure generated by the volatile compo- 65
nents present in the whole crude. In fact, Lutz in U.S.
Pat. No. 4,454,023 (1984) teaches that it is necessary to
pass a whole crude oil through a distillation column
4,778,586
5
tion point. The presence of this amount of oxidant
would not be possible with a liquid which was primarily
carbonaceous.
Other patents which disclose the use of a hydrostatic
column to generate pressure include Beddoes, U.S. Pat. 5
No. 887,506 (1908). Silverman in U.S. Pat. No.
3,371,713 (1968) discloses a method for generating
steam for steam flooding for oil production. Palmer in
U.S. Pat. No. 1,514,098 (1924) discloses a system in
which an elevated vessel is used to provide a low pres- 10
sure hydrostatic head on oil in a thermal cracking vessel.
Other patents include U.S. Pat. No. 3,140,986 of
Hubbard (1964) and U.S. Pat. No. 2,421,528 of Steffen
(1947).
The above-cited patents which disclose vertical tube 15
reactor systems describe the use of such systems with
primarily aqueous streams. None ot these patents describe
treatment of a primarily hydrocarbon stream.
Specifically, there is no suggestion of the thermal treatment
of a hydrocarbon stream in a vertical tube reactor 20
system to provide for viscosity reduction. Based on the
teachings of the visbreaking art as described hereinabove,
it would be expected that coking of the reactor
surfaces would be a significant problem with this configuration.
25
Therefore, it would be advantageous to have a thermal
process by which significant viscosity reduction
can be achieved with a heavy oil feedstock. It would be
particularly advantageous for the process to produce
little or no coke make so that a vertical tube apparatus 30
could be used. Additionally, the process should provide
viscosity reduction without the need for long residence
times and a high throughput rates.
These and other advantages are now achieved by
practice of the present invention as described hereinbe- 35
low.
SUMMARY OF THE INVENTION
It has been discovered that significant improvements
in the transportability of heavy hydrocarbon feeds can 40
result at elevated pressure with the careful control of
the driving temperature differential during relatively
mild thermal treatment of the feed. More particularly,
this invention comprises a method of reducing the viscosity
of hydrocarbon feed comprising: heating said 45
feed at a pressure of at least about 1000 psig to a reaction
temperature of at least about 300° C. by contact with a
heat source; and maintaining the difference between
said reaction temperature and the temperature of said
heat source sufficiently small so as to have minimal coke 50
and enhanced or maximized viscosity reduction at the
reaction temperature and pressure. This is accomplished
by maintaining an efficient heat transfer between
an effiuent product stream and an influent feed
stream in which at least one of the streams is in turbu- 55
lent flow.
This invention further comprises reducing the viscosity
of a hydrocarbon composition by passing a feed
stream of the hydrocarbon composition at an initial
temperature into a vertical tube reactor to form a hy- 60
drostatic pressure head. The influent stream is heated to
a second temperature by heat exchange with an effiuent
product stream in which at least one of the streams is in
turbulent flow. The influent stream is then heated to a
reaction temperature at a reaction pr,essure by contact 65
with an external heat source in which a temperature
differential between the heat source and the hydrocarbon
stream of less than about 30° C. is maintained. The
6
reaction temperature is between about 300° C. and the
coking temperature of the hydrocarbon composition
and the reaction pressure is at least about 1000 psi.
BRIEF DESCRIPTION OF THE DRAWING
The FIGURE is a schematic representation of a preferred
configuration of a vertical tube reactor system
useful in practicing the instant process.
DETAILED DESCRIPTION OF THE
INVENTION
The method of the present invention involves a process
useful for improving crude oil transportability, i.e.,
by treating a whole crude to substantially reduce its
viscosity. In the instant process, a vertical tube reactor
is used to provide the necessary pressure through the
formation of a hydrostatic column of fluid. Coke make
in the reactor is minimized by maintaining a relatively
low driving temperature differential during heating at
the reaction temperature. It has been found that the
necessary reaction temperatures can be attained while
maintaining the low driving temperature differential by
providing substantially improved heat exchange between
the influent feed stream and effiuent product
stream in which at least one of he streams is in turbulent
flow.
As used herein "temperature differential" (AT) refers
to reaction driving force and more particularly, to the
difference between the temperature of the bulk fluid in
the reaction zone (as defined hereinbelow) and the temperature
of the active heat source in a system of indirect
heating. As used herein the "heat transfer surface" refers
to that surface actually contacting the hydrocarbon
stream and providing heat to said stream. The term
"heat source" refers to a heat transfer surface whose
temperature is at least equal to or greater than the temperature
of the hydrocarbon stream which contacts said
surface. The term "active heat source" refers to a heat
source whose temperature is greater than the reaction
temperature but is below the coking temperature of the
hydrocarbon material in contact with the surface.
The temperature differential during practice of the
present invention is minimized to the extent practicable.
It is preferred that the temperature differential be maintained
below about 25° c., more preferably below about
15° C., and most preferably below about 5° C. It has
been found that maintaining a relatively small AT during
treatment of the feed at elevated pressures enables
significantly higher viscosity reductions to be achieved
with minimal or substantially no coke make, e.g. below
about 0.5 weight percent of the hydrocarbon feed, preferably
below about 0.2 weight percent coke make, and
most preferably less than about 0.05 weight percent
coke make. As used herein the term "coke" refers to
material which is insoluble in boiling benzene. As AT
increases, coke make occurs at lower reaction temperatures
and/or at lower pressures and/or at higher final
viscosities, Le. smaller viscosity reductions are achieved
at equivalent coke make.
As used herein the term "reaction temperature"
(TRX) refers to the maximum bulk temperature of the
hydrocarbon stream reached in the process. However,
it is understood that some reaction can begin at a lower
temperature ("initiation temperature"). The maximum
useful temperature in the instant process is the "coking
temperature" of the particular feedstock. The "coking
temperature" is defined herein as the temperature at
which at least about 0.5 weight percent coke is formed
4,778,586
8
It has been found that the reaction zone heat flux required
for practice of this invention is substantially less
than the heat flux required in conventional visbreaking
operations. A typical heat flux for a conventional visbreaker
is ordinarily at least 30,000 BTU/ft2!hour. By
contrast the typical reaction zone heat flux for the
method of the present invention is on the order of about
one-half to less than one-tenth that value or less than
about 15,000 BTU/ft2!hour and more preferably less
10 than about 6,000 BTU/ft2!hour. It is expected that a
heat flux as low as about 2,000 BTU/ft2!hour can be
attained in a commercial scale unit for the present invention.
The pressures useful for the practice of the present
invention are typically above about 1000 psi and preferably
above about 1500 psi in the reaction zone. As used
herein the term "psi" refers to "pounds per square inch
absolute" and "psig" refers to "pounds per square inch
gauge". Such pressures are in excess of those typically
used for visbreaking or most other crude oil treatments
employed at or near the well-site for viscosity reduction
purposes. Similarly, such pressures are in excess of
those used for treating hydrocarbons in the absence of
added hydrogen. Traditionally such high pressures
have been used in conjunction with severe cracking and
thermal treatments where an increase in the hydrogen
to carbon ratio is intended and hydrogenation with
hydrogen gas is most common.
The use of such pressure has an additional advantage
in that the volume percent of the hydrocarbon stream
which is in the liquid phase in the reaction zone is maximized.
This minimizes the concentration of amphaltenes
and other coke precursors and thus reduces the likelihood
of such materials precipitating on internal reactor
surfaces to produce coke.
The process of the present invention is broadly applicable
to reducing the viscosity of petroleum-type hydrocarbons.
The invention is especially useful for treating
heavy oil crudes of a nature and viscosity which
renders them unsuitable for pipeline transport to distant
refineries, i.e. feeds having a viscosity above about 1000
centipoise (cps) at 25· C. (unless otherwise indicated,
viscosity herein is at 25· C.), a pour point above IS· C.
or an API gravity at 25· C. of IS· and below. However,
even "light" heavy crudes, i.e. those having viscosities
of 1000 cps or less, can be beneficially treated as can any
feeds having an API ofless than about 25·. More particularly,
the advantages of reduced viscosity, increased
API gravity and/or reduced pour point can be achieved
by practice of the present invention without regard to
the initial viscosity, API gravity or pour point of the
feed. Additionally, it may be desirable to add a diluent
to the product from the instant process in order to further
reduce the viscosity. It is also possible to blend the
product ofthe instant process with unmodified or virgin
crude to obtain an overall reduction in viscosity of the
final blend product. Heating ofthe product, for example
with heating stations, in order to further reduce the
viscosity or to maintain an acceptable viscosity for a
particular pipeline or transportation medium is also
possible.
Heavy hydrocarbon feeds to the· process of the instant
invention comprise, but are not limited to, heavy
whole crude oil, tar sands, bitumen, kerogen, and shale
oils. Examples of heavy crude oil are Venezuelan Boscan
crude oil, Canadian Cold Lake crude oil, Venezue-
Ian Cerro Negro crude oil and California Huntington
Beach crude oil. The viscosity ofthe typical feed at 25·
7
based upon the hydrocarbon feed. In ordinary operation,
the reaction temperature is maintained below the
coking temperature. At a minimum the reaction temperature
used for practice of the present invention is high
enough to initiate a thermal cracking reaction at an 5
effective rate. For most feeds the reaction temperature
is above about 300· C. and less than about 475· C., more
typically in the range of about 350· C. to about 450· C.
and most often in the range of about 375· C. to about
435· C.
The influent hydrocarbon stream is introduced to the
inlet of the vertical tube reactor at a first or initial temperature
(TI), normally less than about 100· C., and an
initial pressure (PI) typically less about 200 psi. As any
particular volume element of the influent hydrocarbon 15
stream travels down the downcomer in the vertical tube
reactor, the pressure on the increment increases due to
the increasing hydrostatic column of fluid above it.
Additionally, the bulk of the influent stream increases to
a second temperature (T2) due to heat exchange with 20
the effluent product stream. The second temperature is
the highest bulk temperature reached in the influent
stream due to heat exchange with the effluent stream.
Normally this temperature is at least about 200· C.,
preferably this temperature is at least about 250· C., and 25
preferably this temperature is at least about 300·. In the
reaction zone, the temperature of the hydrocarbon is
increased to a maximum reaction temperature (TRX)
due to contact with an active heat source. As used
herein, the term "reaction zone" refers to the region in 30
the vertical tube reactor in which the bulk temperature
of the hydrocarbon stream is greater than the second
temperature (T2) and equal to or less than the reaction
temperature (TRX). This temperature is achieved by
contacting the hydrocarbon stream with the active heat 35
source.
In order to minimize the temperature differential, the
second temperature T2 should be maximized. Therefore,
it is necessary for the heat exchange between the
influent and effluent streams to be more efficient than 40
those disclosed in the known patents relating to vertical
tube reactors. The temperature of the influent stream
achieveable by heat exchange with the reaction product
is limited by a number of factors including the temperature
of the reaction product, the heat-exchange surface 45
area and the velocity of the hydrocarbon streams. In
order to achieve the necessary heat-exchange efficiencies,
it has been found that turbulent flow of the streams
is necessary. Although static mixing devices can be used
to provide turbulent flow, this is not preferred. It has 50
been found that substantially improved results are obtained
when at least one of, and preferably both, the
influent feed stream and the product stream are in substantially
vertical, multiphase flow. When both streams
are in vertical multiphase flow, an increase in heat- 55
exchange efficiency of at least about 100% can be
achieved compared to heat exchange when neither
stream is in turbulent This allows a T2temperature to be
attained which is sufficiently close to the reaction temperature
to allow a small A.T to be used in order to 60
provide the incremental heat necessary to attain the
desired reaction temperature.
It has been found that thermal treatment of hydrocarbon
feeds according to the present invention, wherein
A.T is minimized, results in advantageous viscosity re- 65
ductions with significantly less heat flux in the reaction
zone. Heat flux is defined herein as the heat flow (Q)
into the feed fluid per unit area of heat transfer surface.
4,778,586
9
c. can vary widely ranging from about 300,000 cps or
more to about 20,000 cps or lower. In practice, as would
be expected, the most significant reductions in viscosity
are achieved where the starting feed is most viscous. It
has been found that essentially unpumpable feeds hav- 5
ing viscosities up to about 200,000 cps can be rendered
suitable for pipeline transport by treatment according to
the present invention. With feeds of viscosities greater
than about 200,000 cps, significant viscosity reduction,
preferably greater than 50 percent, more preferably 10
greater than 90 percent, and most preferably greater
than 95 percent (based on feed viscosity) is achieved by
the method of the present invention, although supplemental
physical treatment, such as heating or dilution,
can still be used to render the product more readily 15
pumpable.
In a similar manner, the process of the present invention
is effective to reduce the pour point and/or increase
the API gravity of the feed. Typically, a reduction
ofat least about 15" C. in pour point is preferred. In 20
particular, for feeds having a pour point of between
about 15" C. and about 30" C., the process of the present
invention can yield a product with a pour point below
about -10" C. For typical heavy feeds having an API
gravity of less than about 25" and more typically less 25
than about 15", the process of the present invention can
yield a product with an API gravity increase of at least
about 2", .
Typically, the feeds to the process of the present
invention are whole crudes, "untopped", i.e. without 30
passing through a distillation unit to remove lower boiling
components, and without added solvents. However,
the advantageous results of the present invention can be
achieved with separate crude fractions and independent
of any solvents or water which are present. Ordinarily, 35
whole crude contains water with the amount of water
depending upon the method of production. Crude oil
produced by steam flood commonly contains in excess
of 50 weight percent water as measured at the wellhead.
It is contemplated that the feedstock for the instant 40
process normally passes through the usual primary
water/oil hot phase separator to remove most of the
aqueous phase and reduce the water level to less than
about 10 weight percent and preferably less than about
5 weight percent of the hydrocarbon feedstock. The 45
terms "hydrocarbon stream", "hydrocarbon feedstock",
and "hydrocarbon feed" are used interchangeably
herein to mean the fluid stream which is passed
through the instant process and contains primarily hydrocarbonaceous
components but can also contain 50
smaller amounts of other components such as water.
As expected, treatment by heating, according to the
present invention, results in some conversion or alteration
of the hydrocarbon feed. However, it has been
found that even at constant conversion percentages, (Le. 55
conversion of the +950" F. fraction), use of elevated
pressure according to the present invention results in
enhanced viscosity reduction.
It is generally known that increased temperature in
the thermal treatment of hydrocarbons results in de- 60
creased viscosity due to higher conversion, Le. increased
formation of lighter products, and a concomitant
increase in coke formation. Avoidance of coke
formation by use of more moderate temperatures in
visbreaking processes, heretofore has required unduly 65
long "soaking" or residence times on the order of 2-24
hours to effect any significant results. Surprisingly, it
has been found that temperatures high enough to effect
10
significant viscosity reduction can be used without
causing significant coke make and/or without the need
for long residence times by the use of elevated pressure
and a minimal temperature differential. Reaction and/or
residence times in the reaction zone for processes of the
present invention are relatively short, Le. times less than
1 hours, often less than 30 minutes, more frequently less
than about 15 minutes and even less than about 5 minutes
are possible.
Heretofore, the relationships between reaction temperature,
AT, pressure and coke make as they specifically
relate to viscosity reduction have gone unrecognized.
Practice of the processes of the present invention
permits valuable viscosity reduction to be maximized at
elevated pressures above 1,000 psi by use of a reactor
temperature and a related AT selected to minimize coke
make. By the processes disclosed herein, it becomes
possible to maximize viscosity reduction under practical
conditions of minimal coke make and relatively low
temperatures by using high pressures, e.g., greater than
1,000 psi, and minimizing the system AT. While it is
anticipated that in normal operations the primary objective
is to maximize viscosity reduction, it is recognized
that particular circumstances may require a different
mode of operation whereby somewhat less than the
absolute "maximum" viscosity reduction results. For
example, if heating stations and insulated pipelines are
available, it may be desirable to increase throughput and
accept a smaller reduction in viscosity. As will be understood
by those skilled in the art the terms "maximize"
or "maximizing" and "minimum" or "minimizing"
are not absolute and are intended to encompass
selection of parameters which approach such maximums
or minimums.
The use of a vertical tube reactor involves subjecting
a moving hydrocarbon feed stream to essentially continually
increasing pressure until a reaction pressure (P2) is
reached. As used herein the term "reaction pressure"
refers to the maximum pressure on the hydrocarbon
stream in the reaction zone. The hydrocarbon stream is
maintained at a reaction temperature of about 300" C. to
about 475" C., more commonly about 350" C. to about
450" C. and a reaction pressure of at least about 1000 psi
for a time sufficient to provide the desired reduction in
viscosity of the hydrocarbon stream. As used herein the
terni"treated hydrocarbon stream" refers to the product
Of the instant process in which the viscosity of the
hydrocarbon stream has been reduced without significant
coke make. It is preferred that the pressure of the
resulting treated hydrocarbon stream is essentially continually
decreased to an exit pressure (P3).
The temperature of the hydrocarbon stream is also
essentially continually increased from an initial temperature
to a second temperature by heat exchange with
the treated hydrocarbon stream. The bulk temperature
of the stream is then increased to a reaction temperature
by contact of the stream with an active heat source. The
temperature of the resulting treated hydrocarbon
stream is essentially continually decreased from the
reaction temperature to a fmal temperature by heat
exchange with influent feed stream.
The hydrocarbon stream is ordinarily a whole crude
oil which has been subjected to the primary dewatering
process discussed hereinabove. However, it is contemplated
that any of the other heavy hydrocarbon streams
discussed hereinabove such as bitumen, shale oil or resid
could be subjected to this embodiment of the instant
process. If the hydrocarbon stream is whole crude, the
4,778,586
11 12
initial temperature of the incoming stream is ordinarily and, in fact, a small vapor phase can be beneficial in
about 40· C. to about 100· C. depending upon the promoting mixing of the stream for rapid distribution of
method of production. In general, the present invention heat. Preferably, the vapor phase should amount to no
is operable independent of the presence or absence of more than about 10 volume percent of the hydrocarbon
water in varying amounts. 5 stream and preferably less than 5 volume percent. If the
The pressure on any particular volume segment of vapor phase comprises a substantial percent of the
the hydrocarbon stream is essentially continuously in- stream volume, it can become difficult to maintain a
creased from an initial pressure to the reaction pressure. pressure balance in the reactor vessel.
By "essentially continuously" it is meant that the stream Preferably, the temperature of the incoming hydrois
not maintained at a constant pressure below the reac- 10 carbon stream is increased essentially continuously
tion pressure for a significant period of time, i.e. any from an initial temperature to the second temperature
period of constant pressure that has a duration of less T2. By "essentially continuously" it is meant that there
than about 5 minutes and ordinarily less than about one are no long soaking periods in which the stream is mainminute.
It is possible that phase changes can occur de- tained at a constant temperature. During this temperapending
upon the composition of the stream. This can 15 ture increase, it is possible for various phase changes to
result in rapid pressure increases or decreases possibly occur in the stream. For example, depending upon the
followed by momentary leveling of pressure. However, temperature and pressure, water contained in the stream
except for such stream composition-dependent devia- can vaporize. Such phase changes can cause a tempotions,
the increase in pressure is continuous from the rary leveling or even a decrease in the temperature of
initial pressure to about the reaction pressure. 20 the stream due to the heat of vaporization. However,
In operation of the instant process, the pressure on such a leveling or dip in temperature is of short duration
the stream ordinarily increases from some lower pres- and in the instant process the temperature increase
sure, when the bulk temperature of the stream is at the quickly resumes.
second temperature, to the reaction pressure as the The temperature of the influent hydrocarbon feed
stream passes through the reaction zone. This operation 25 stream is increased by contact with a heat source. The
contemplates that the flow of the stream through the heat source can be any means capable of providing the
reaction zone is substantially linear or plug flow. If necessary temperature increase in the hydrocarbon feed
another manner of flow through the reaction zone is stream from the initial temperature to the second temused,
e.g. if there is substantial backmixing of the perature T2. For example, multiple zones of increasing
stream, it is possible that a particular segment of the 30 temperature can be provided by electrical resistance
stream would be exposed to some fluctuation in pres- heaters or through use of a heat exchange fluid. The
sure. However, the maximum range ofany such fluctua- heat source should be maintained at a temperature
tions is expected to be from between the pressure at the below the reaction temperature in order to assure minisecond
temperature and the reaction pressure. As set mum coke make. The influent and effluent hydrocarbon
forth hereinabove, the reaction pressure is at least about 35 streams should be in thermal communication with one
1000 psi and preferably at least about 1500 psi. In nor- another to provide for maximum efficiency. Economimal
operation it is not expected that the reaction pres- cally it is preferred that the influent and effluent streams
sure would exceed about 4000 psi. Commonly, the reac- be in counter-current heat exchange in which the
tion pressure ranges from about 1000 psi to about 3000 treated hydrocarbon stream is initially contacted at its
psi and usually ranges from about 1000 psi to about 2000 40 highest temperature with the influent hydrocarbon feed
psi. The initial pressure of the hydrocarbon feed stream stream at or near the reaction zone. The effluent prodis
ordinarily between about 25 psi and about 1000 psi, uct fluid is then maintained in countercurrent heat exand
preferably is between about 25 psi and about 500 change contact with the influent hydrocarbon stream to
psi. It is contemplated, however, that the hydrocarbon provide an essentially continuous increase in the temfeed
stream can be provided under a higher initial pres- 45 perature of the influent stream and a continuous desure
if it is desired to have a higher reaction pressure crease in the temperature of the effluent fluid. Other
than is obtained by the hydrostatic head of the fluid things being equal, it is anticipated that the time recolumn.
As set forth hereinabove, the reaction pressure quired to heat the influent hydrocarbon feed from its
is primarily due to a hydrostatic head. If it is desired initial temperature to a second temperature (heat exthat
the reaction pressure be greater than would be 50 change temperature) is at least about 30 seconds and
generated by the hydrostatic head, the initial pressure of preferably at least about 100 seconds.
the hydrocarbon feed stream can be increased by, for In normal operation the hydrocarbon feed stream is
example, centrifugal pumps to provide the desired total heated to the second temperature which is preferably
reaction pressure. within about 30· C. of the reaction temperature before it
The high pressure serves to maintain in liquid phase 55 contacts an "active heat source". As discussed hereinvolatile
components present in the hydrocarbon feed above, the differential between the temperature of the
stream or formed during thermal cracking reactions. bulk hydrocarbon fluid at reaction temperature and the
While the process of coking is not fully understood, it is active heat source should be maintained as low as possiknown
that materials such as asphaltenes are more ble, normally below about 30· C., preferably below
likely to form coke. Once these materials precipitate 60 about 25· C., more preferably below about 15· C., and
and solidify on surfaces it is difficult to dissolve them most preferably below about 5· C. In addition to minibefore
coke deposits are formed. It is therefore impor- mizing actual coke make, this IlT provides a product
tant to maximize the liquid phase in the reaction zone to which has good stability in storage and during transporminimize
the concentration of asphaltenes and other tation, Le. solid materials do not form and precipitate.
coke precursors to avoid the precipitation from the 65 Ordinarily the reaction temperature for a whole
hydrocarbon phase and possible deposition on internal crude oil feedstock is in the range of about 300· C. to
reactiQn surfaces with subsequent coke formation. A about 450· C. and preferably between about 375· C. and
small volume fraction of the stream can be vapor phase about 435· C. The hydrocarbon stream is maintained at
4,778,586
13
the reaction temperature and pressure for a time sufficient
to effect the desired viscosity reduction without
providing significant coke make. In normal operation,
the hydrocarbon stream is maintained at the reaction
temperature for less than 1 hour, preferably less than 30 5
minutes, and most preferably less than 15 minutes. Ordinarily
the viscosity of the treated or modified stream is
reduced by at least SO percent and usually by at least 90
percent and more preferably by at least 95 percent compared
to the untreated feedstock. 10
This treated hydrocarbon stream is passed out of
contact with the active heat source. The temperature
and pressure of the treated stream are reduced essentially
continuously from the reaction temperature and
pressure to a final or exit temperature (TE) and pressure 15
P3 by heat exchange contact with the feed stream.
While the temperature and pressure are being reduced,
phase changes can occur, for example, water vapor can
condense to form liquid water. This can result in a momentary
leveling in temperature due to the latent heat 20
of vaporization. Also the pressure can rapidly drop due
to this condensation. These are transient phenomena
dependent upon the particular composition of the
stream. Therefore, when the temperature and pressure
changes are viewed as a whole, the decreases are essen- 25
tially continuous from the reaction conditions to the
final conditions.
Although some pressure reduction occurs as the result
of a reduction in temperature, there is a continual
reduction in pressure as the hydrostatic pressure head is 30
decreased.
The use of a hydrostatic pressure head is particularly
useful when whole crude oils or other feedstocks which
contain a substantial amount of volatile components,
e.g. materials boiling below about 300· C. This is even 35
more critical when the feedstock contains a significant
amount of water. These materials are not readily useable
in conventional visbreaking processes due to the
high pressures required in order to provide an accept-,
able residence time at reaction temperature. In the in-. 40
stant process, the necessary pressures can be provided
with simple, relatively inexpensive equipment.·· .
It is particularly important in a vertical tube reactor
for the coke make to be minimized in the process. Excessive
coke formation can rapidly coat the internal 45
surfaces of the apparatus and cause premature shutdowns.
Therefore, the coke make should be kept below
about 0.5 weight percent and preferably below about
0.2 weight percent. As discussed hereinabove this is
accomplished by a combination of very efficient heat 50
exchange between the influent and effluent streams and
a low ~T in the reaction zone.
The exit temperature and pressure depend on the
feedstock being used, the particular reaction conditions,
and the extent of viscosity reduction desired in the 55
feedstock. Ordinarily, the temperature ranges from
about 75" C. to about 200· C. and the pressure ranges
from about 150 psig to about 350 psig.
The instant invention can be more readily understood
after a brief description of a typical application. As will 60
be understood by those skilled in the art, other apparatus
and configurations can be used in the practice of the
present invention.
The FIGURE depicts a subterranean vertical reactor
10 disposed in a well bore 12. The term "vertical" is 65
used herein to mean that the tubular reactor is disposed
toward the earth's center. It is contemplated that the
tubular reactor can be oriented several degrees from
14
true vertical, i.e. normally within about 10 degrees.
During operation, flow of the hydrocarbon stream can
be in either direction. As depicted, flow of the untreated
hydrocarbon feed stream is through line 13 and into
downcomer 14 to the reaction zone 16 and up the concentric
riser 18. This arrangement provides for heat
exchange between the outgoing product stream and the
incoming feed stream.
During start-up, untreated hydrocarbon feed is introduced
into the vertical tube reactor system through feed
inlet 13, the flow rate being controlled by valve 20. The
hydrocarbon feed stream passes through downcomer 14
into reaction zone 16 and up through concentric riser 18
exiting through discharge line 22. Unless external heat is
provided to the hydrocarbon stream, the initial temperature
TI is equal to the final heat exchange temperature
T2 and is also equal to the maximum temperature in the
reaction zone TRX(provided there is no heat loss to the
environment). It is necessary to increase the temperature
of the effluent stream so that the desired T2 temperature
of the influent stream can be obtained. This can be
accomplished by passing the influent stream through an
above-ground heating means 24 so that the T I is essentially
equal to the desired T2. Alternatively, the necessary
heat can be provided by an external heating means
26 surrounding the reaction zone. In another configuration
(not shown), the downcomer 14 can be jacketed to
allow external heating of the hydrocarbon stream at this
location in addition to or instead of heating at the reaction
zone. Of course, the external heating means 26, can
be used in conjunction with the above-ground heating
means 24 to provide the hydrocarbon feed stream at the
desired temperature T2. It may be necessary during
start-up to provide a hydrocarbon feed stream which
has a lower viscosity than the hydrocarbon material to
be processed during normal operation to allow ready
transport of the fluid through the reactor system. Additionally,
it is preferred during start-up operation for the
effluent stream to be recycled by diverting through
valve 28 into recycle line 30. This recycle allows conservation
of energy necessary to heat the hydrocarbon
stream and the apparatus to the desired T2 temperature.
Once the desired T2has been attained, temperature of
the exterIlal heating means 26 can be increased to provide
the desired TRX in the reaction zone. Recycle
through line 30 can be stopped and the feed which is
desired to be processed can be directed into the vertical
tube reactor through line 13. As the treated hydrocarbon
exits the vertical tube reactor through line 22, it can
be directed to an above-ground product treatment
means 32 which can separate gaseous materials such as
methane from the product stream. A fraction of components
boiling below about 40· C. can also be separated
and recycled into the feed stream through line 34. As is
discussed in more detail hereinbelow, the recycle of
such volatile materials, such as butanes and pentanes
can be used to induce multiphase flow in the downcomer
14 to provide for significantly improved heat
exchange.
As the influent hydrocarbon stream passes down
through downcomer 14, any particular volume segment
is exposed to increasing pressure due to the hydrostatic
column of fluid above it. The temperature of the hydrocarbon
stream is measured by temperature monitors 36
which can be located in the hydrocarbon stream
throughout the vertical tube reactor system as desired.
Pressure monitors 38 can also be located throughout the
4,778,586
EXAMPLE 1
The batch autoclave and the continuous flow unit
experiments described above were performed on the
Cold Lake crude oil samples. Analysis of the feed for
these tests is given in Table lAo Results from a mass
spectrometer analysis of the 273° F.-430° F. fraction of
the Cold Lake feed are given in Table lB.
The experimental conditions and analysis of the products
are given in Table IC.
16
Cold Lake heavy crude oil and on the four Venezuelan
crudes.
The batch experiments were performed in rocking
bomb autoclave units. The continuous-flow bench unit
experiments were performed in a specially designed
system, containing the following sections: a high pressure
feed system, a tubular reactor, and a pressure letdown
system. The unit was designed to handle flow
rates of 0.2 to 2.2 gallonslhr. at temperatures up to 450'
C. and pressures of 3000 psi. The feed system consisted
of an electrically heated five gallon tank connected to a
recirculation pump. The heavy oil feed was recirculated
continuously through in-line heaters and back into the
tank to keep the oil well mixed and to maintain the oil
temperature at 70° C. A side stream from the recirculation
system served as the feed to the tubular reactor
through a high pressure system pump. An additional
three gallon heated tank supplied a high temperature oil
to the system for start up and shut down. The reactor
consisted of 50 feet of i inch O.D. stainless steel tubing
coiled to form a 9-inch diameter coil with 2-inch spacing
between each ring of the coil. Reaction temperature
was reached and maintained by means of a fluid bed
sand bath. Temperature was measured throughout the
system including two points within the heated coil section.
The coil form, coupled with the uniformity of the
heated fluidized sand bed, allowed a fine degree of
temperature control with temperature differences between
the sand bed and the oil of less than 5° C. Pressures
were measured at various points in the circuit.
The temperature and pressure of the oil was measured
as it exited from the tubular reactor. The pressure of the
product was decreased through a series of valves, and
the product was collected in a low pressure receiver
tank. In the low pressure receiver tank, the liquid and
gas phases separated, with the liquid exiting the bottom
and gas sampling and venting at the top.
For each experiment, the products were analyzed for
water content, viscosity, density, distillation fractions,
solids content, asphaltenes content, Conradson carbon,
sulfur content and gas composition. Additionally, tests
were made with feed containing added water of approximately
2 percent, 5 percent, and 10 percent by weight
to determine the effects of water on the products and on
process parameters. The runs with added water are tests
CBU-9 to -11, -19 to -21, and -23 to -25.
The products from the batch and the continuous-flow
50 tests were analyzed for structural components and compared
with the structural components of the crude oil
feed. The structural data were obtained by mass spectral
analysis. The structural data on the crude oil feeds
were determined by analysis of whole oil samples. The
structural data on the products were determined by
separate analysis on distillation cuts of the product. The
result for the whole oil product was then calculated
from these results.
EXPERIMENTAL
15
vertical tube reactor system to monitor any pressure
increases or fluctuations in the fluid stream.
The external heat 26 source preferably uses a heat
exchange fluid which is passed into inlet 40 through a
jacket surrounding the reaction zone and out through 5
outlet 42. The use of the heat exchange fluid allows
careful temperature control to assure that the desired
temperature differential can be maintained. Additionally,
control of the heat exchange temperature can assure
that the surface temperature of the vertical tube 10
reactor in the reaction zone does not exceed the coking
temperature.
In order to obtain the desired T2 temperature of the
influent stream by heat exchange with the effluent
stream, it is necessary that very efficient heat exchange 15
be provided. It has been found that unexpectedly higher
overall heat transfer coefficients than would be predicted
from empirical heat transfer correlations such as
Sieder-Tate can be attained by providing substantially
vertical, multiphase flow in the fluid stream. If neces- 20
sary, multiphase flow can be induced in the influent
stream by recycling volatile components from the effluent
product stream to provide a gas phase in the liquid
phase. As the influent stream progresses down down- 25
comer 14, the increasing pressure serves to liquify and/
or dissolve the gaseous components in the liquid phase
providing for substantially a liquid phase in the reaction
zone. The substantially liquid phase in the reaction zone
is desired in order to minimize the concentration of 30
asphaltenes and other coke producing materials in the
reaction zone in order to minimize coke formation on
surfaces in the reaction zone. As the effluent product
flows up the riser, the pressure on any particular volume
segment decreases. Volatile components dissolved 35
in the liquid at reaction pressure can vaporize to yield a
vapor phase in the liquid stream and provide multiphase
flow in the effluent stream. The efficient heat exchange
allows the heat flux required in the reaction zone to be
minimized. Thus, the typical heat flux in the reaction 40
zone is substantially less than that required in an conventional
visbreaker operation. To maximize heat exchange
efficiency, it is preferred that both the influent
and effluent streams be in multiphase flow, although
improved efficiency can be obtained if only one of the 45
streams is in multiphase flow.
The following examples are intended by way of illustration
and not by way of limitation.
In the following examples, five heavy crude oils and
two shale oils were used to test various process parameters.
One of the crude oils came from Cold Lake, Alberta,
Canada and four of the crudes came from Venezuela.
The Boscan and Tia Juana crudes were from the 55
Lake Maracaibo Basin and the Zuata and Cerro Negro
heavy oils were from the Orinoco River area. In addition,
heavy shale oils were tested.
The heavy crude oils and shale oil were analyzed for
water content, viscosity, density, distillation fractions, 60
solids content, asphaltenes content, pour point, Conradson
carbon, and sulfur content. Additionally, the pour
point and the salt content, as chloride, was measured for
the Venezuelan heavy oils.
In order to test the different parameters for heavy oil 65
conversion, including the effect of temperature, pressure,
residence time, and water content of the feed oils,
both batch and continuous-flow testing was done on the
17
TABLEIB
MASS SPECTROMETER ANALYSIS OF
285-430' F. FRACTION OF THE COLD LAKE FEED
Paraffins 35.3 vol %
Olefins NO
4,778,586
5
18
TABLE IB-continued
MASS SPECTROMETER ANALYSIS OF
285-430' F. FRACTION OF THE COLD LAKE FEED
Cycloparaffins 35.0
Condo Cycloparaffins' 29.0
Alkyl Benzenes ~
100.0 vol 7%
"May include cyclic olefins and certain sulfur compounds.
ND None detected.
10
TABLEIA
Temp. Range. 'F. at I Atmos.
Cut Vol % of Whole Oil
I Vol. % OH at Cut End
Cut Wt % of Whole Oil
I Wt % OH at Cut End
'API Gravity 60/60
Specific Gravity 60/60
Sulfur, wt %
Nitrogen, wt %
Pour Point, 'F.
Cetane Index(2)
Smoke Point, mm
Can Carbon Res, wt %
Viscosity,
100' F., cst
210' F., cst
275' F., cp
Nickel, wppm
Vanadium, wppm
ANALYSES ON COLD LAKE CRUDE
Whole Oil IPB-285 285-430 430-525
100 No 0.99 3.05
100 Material 0.99 4.04
100 0.83 2.67
100 0.83 3.50
10.4 36.9 30.4
0.9969 0.8402 0.8742
4.44 1.06 1.30
<-75
35.4
11.1
3.02
1.19
525-650
1\.16
15.20
10.10
13.60
25.4
0.9017
1.94
122 ppm
-75
39.5
9.9
6.01
1.69
650-950
34.24
49.44
32.86
46.46
16.4
0.9567
3.31
0.14
5
25.2
(3)
0.39
149
9.34
7.4
ND
950+
50.56(1)
100.00(1)
53.54(1)
100.00(1)
2.8
1.0539
5.91
24.4
2,930
131
284
Sulfur balance closure = 101.2%.
ND = None Detected. .
(llBy difference to give 100% recovery since loss is primarily in the residue.
(2)Calculated from midpoint of distillation fractions. not from a separate D-86 distillation.
(31Material would not wick, test not applicable.
35
40
45
50
55
60
65
TABLEIC
COLO LAKE HEAVY OILS RUN OATA
Pres· Feed Product Viscosity" Residual Asphaltene* Solid Coke Gas IPB· 450- Resid Con· Sulfur
Temp sure, H2O Time H2O cp cp Gravity WI. Conv. WI. Alter. WI. WI. Wt. 450' F. 950' F. +950 F. Carbon WI.
Run 'c. psig % min··· % 25' C. 80' C. 'API % ,% % % % % % WI. % WI. % WI. % WI. % %*
Cold Lake Crude (Barrel I) - Batch Tests
Feed 0.7 41,600 687 11.5 60.2 16.3 0.00 0.05 4.7 35.1 60.2 I\.7 4.5
Run I 360 290 0.7 15 Trace 26,400 550 12.2 58.4 3.0 16.1 \.2 0.00 0.0 0.3 2.0 39.3 58.4 10.8 4.5
Run 2 380 330 0.7 15 0.5 9,710 334 13.2 56.0 7.0 14.2 12.9 0.00 0.0 0.3 4.7 39.0 56.0 I \.9 4.6
Cold Lake Crude (Barrel 2) - Batch Tests ... Feed 0.2 47,100 886 I\.4 59.0 16.3 0.00 0.2 3.9 36.9 59.0 4.6 \C
Run I 370 250 0.2 15 Trace 16,300 370 12.9 56.1 4.9 14.5 11.0 0.00 0.0 0.3 6.4 37.2 56.1 4.5
Run 2 415 710 0.2 15 0.0 156 27 18.6 37.2 37.0 13.2 19.0 0.14 1.5 I.3 11.7 48.3 37.2 12.9 3.9
Run 3 405 340 0.2 15 Trace 758 58 13.7 45.7 22.5 14.0 14.1 0.00 0.0 1.I 8.3 44.9 45.7 4.1
Continuous Unit Runs (Barrel 2)
CBU·I 400 40 0.2 1.8 15,400 327 14.2 12.9 0.00
400 40 0.2 2.2 0.0 13,900 333 13.6 56.8 3.7 14.2 12.9 0.00 NO 0.4 4.1 38.7 56.8 11.6 4.5
415 20 0.2 0.6 0.0 10,400 347 13.9 5\.9 12.0 14.2 12.9 0.00 NO 0.5 5.9 4\.7 51.9 10.8
415 20 0.2 0.6 0.0 8,300 243 13.1 53.5 9.3 13.4 17.8 0.00 NO 0.5 5.3 40.7 53.5 4.3
CBU-2 400 390 0.2 2.4 4,810 177 13.3 18.4 0.00
400 400 0.2 2.7 Trace 4,080 148 12.1 53.6 9.2 13.1 19.6 0.00 NO 0.9 6.9 38.6 53.6 12.2 4.5
400 1040 0.2 4.6 2,470 111 12.5 23.3 0.03 4.4 ~
400 1060 0.2 2.8 Trace 2,810 126 12.7 49.3 16.4 12.9 20.9 0.00 NO I.3 9.7 39.7 49.3 12.4 4.3 ':...1
415 910 0.2 3.5 664 61 13.3 18.4 0.00 .......
415 920 0.2 3.5 Trace 506 43 14.5 43.5 26.3 13.2 19.0 0.00 NO 2.9 8.9 44.7 43.5 13.0 4.2 20
415 390 0.2 2.6 819 64 13.1 19.6 0.02 VI
00
415 430 0.2 2.6 Trace 776 64 12.5 47.1 20.2 13.2 19.0 om NO 3.1 9.7 40.1 47.1 12.3 4.3 ~
CBU-3 415 1000 0.2 3.4 Trace 723 54 13.0 45.4 23.1 13.5 17.2 0.00 NO 2.2 9.2 43.2 45.4 12.4 4.1
425 990 0.2 2.9 Trace 281 29 13.6 39.2 33.6 13.5 17.2 0.00 NO 3.2 11.5 46.6 39.2 13.5 4.0
435 1020 0.2 2.3 Trace 175 23 14.5 37.2 36.9 13.7 16.0 0.04 NO 4.2 12.9 45.7 37.2 13.3 4.0
445 1020 0.2 2.0 Trace 63 9 16.7 29.3 50.3 12.7 22.1 0.06 NO 5.8 18.1 46.8 29.3 12.5 3.8
CBUA 415 2010 0.2 5.4 Trace 435 40 13.5 45.0 23.7 13.3 18.4 0.00 NO 4.3 8.1 42.6 45.0 10.6 4.0
425 2060 0.2 4.5 0.0 245 25 14.5 39.3 33.4 13.2 19.0 0.02 NO 5.9 9.4 45.4 39.3 12.8 3.9
435 2020 0.2 5.4 Trace 52 16 16.0 31.9 45.9 11.5 29.4 0.17 NO 6.3 14.4 47.4 3\.9 11.4 3.7
445 2020 0.2 3.5 Trace 25 9 16.8 28.5 51.7 9.4 42.3 0.00 NO 8.2 16.1 47.2 28.5 II.l 3.7
CBU-5 435 1010 0.2 8.3 0.0 85 20 14.2 33.3 43.6 14.5 I \.0 0.23 NO 7.4 11.5 47.8 33.3 12.7 4.0
CBU-6 415 1060 0.2 4.1 0.0 442 49 13.9 45.5 22.9 13.0 20.3 0.00 NO 4.2 6.3 44.0 45.5 12.4 4.2
415 1010 0.2 2.7 0.0 1,250 74 13.9 48.7 17.5 12.8 2\.5 0.00 NO 2.3 6.1 42.9 48.7 12.6 4.3
425 940 0.2 4.3 0.0 219 26 14.4 44.8 24.1 13.3 18.4 0.00 NO 3.7 6.3 45.2 44.8 13.3 4.1 N
425 1030 0.2 2.7 0.0 605 46 14.2 47.1 20.2 13.0 20.3 0.00 NO 2.5 5.8 44.6 47.1 12.2 4.3 0
CBU-8 425 1040 0.2 2.6 0.1 259 33 14.1 37.8 35.9 12.8 2\.5 0.13 NO 4.8 15.5 4\.9 37.8 12.5 4.2
425 1030 0.2 2.6 0.1 841 68 13.2 43.1 27.0 12.9 20.9 0.08 NO 1.6 13.3 42.0 43.1 12.4 4.3
435 1060 0.2 2.3 0.1 163 22 14.2 35.3 40.2 13.3 18.4 0.14 NO 4.7 18.0 42.0 35.3 13.2 4.0
435 1000 0.2 2.2 0.05 222 27 13.9 39.1 33.7 13.7 16.0 0.17 NO 4.1 17.5 39.4 39.1 13.6 4.2
445 1020 0.2 2.5 0.05 69 9 15.3 29.2 50.5 11.5 29.5 0.08 NO 5.3 24.2 4\.4 29.2 12.9 4.0
445 1010 0.2 \.7 0.05 198 26 13.9 37.2 37.0 13.5 17.2 0.21 NO 4.6 20.9 37.4 37.2 13.7 4.2
CBU-9 5.1 39,300 1090
Feed 415 1150 5.1 1.9 4.4 3,300 227 12.5 53.9 8.6 13.2 19.1 0.11 NO 2.9 2.6 40.6 53.8 12.4 4.4
415 2080 5.1 3.1 4.6 1,730 141 12.6 52.8 10.5 13.0 20.3 0.07 ND 4.5 2.6 40.2 52.8 11.8 4.4
425 1040 5.1 1.5 4.8 1,280 84 14.0 53.9 3.6 13.3 18.4 0.08 NO 3.7 2.7 39.8 53.9 12.6 4.3
425 2020 5.1 3.1 3.1 1,100 86 13.9 48.0 18.6 12.8 2\.5 0.05 NO 3.9 4.9 43.1 48.0 12.1 4.2
CBU-IO 5.1 45,600 816
TABLE Ie-continued
COLO LAKE HEAVY OILS RUN OATA
Feed 435 1040 5.1 1.4 3.2 572 52 13.2 46.7 20.6 13.2 18.9 0.08 NO 4.9 5.8 42.6 46.7 12.9 4.1
435 2050 5.1 3.0 3.4 372 44 13.3 44.5 24.6 13.1 19.6 0.15 NO 3.7 6.5 45.3 44.5 13.1 4.2
445 1070 5.1 1.8 2.0 283 35 13.9 40.6 ~1.2 14.3 12.4 0.11 NO 5.7 8.6 45.1 40.6 13.8 4.2
445 2050 5.1 3.2 0.0 110 18 15.0 36.1 38.8 12.5 23.3 0.14 NO 6.4 11.3 46.2 36.1 12.2 3.8
CBU-II 10.7 42,300 1,060
Feed 415 2060 10.7 3.2 7.5 3,730 132 13.0 50.6 14.2 13.4 17.8 0.10 NO 2.1 5.3 42.0 50.6 12.3 4.2
425 2070 10.7 2.9 6.4 1,300 73 13.6 46.2 21.7 12.8 21.4 0.12 NO 2.9 7.8 43.0 46.2 12.4 4.3
435 2050 10.7 2.6 7.5 510 44 14.1 41.3 30.0 14.4 11.8 0.21 NO 3.7 11.5 43.6 41.3 13.1 4.0
445 2030 10.7 2.5 4.7 260 41 16.0 38.8 34.2 15.9 2.1 0.37 NO 6.7 7.5 46.9 38.8 13.8 4.1
CBU·12 415 1030 0.2 7.1 0.0 480 37 13.2 43.4 26.4 13.5 17.2 0.11 NO 3.8 8.0 44.8 43.4 13.3 4.3 N
425 1040 0.2 5.6 0.1 248 26 13.8 37.6 36.3 13.8 15.5 0.23 NO 3.8 13.2 45.4 37.6 13.4 4.1
~
435 1050 0.2 4.9 0.1 53 15 16.0 30.8 47.8 11.1 32.1 0.05 NO 5.7 18.3 45.2 30.8 11.9 4.0
445 1080 0.2 3.0 0.0 20 12 17.4 24.6 58.3 9.7 40.7 0.08 NO 12.4 19.8 43.2 24.6 11.4 3.9
CBU-13 445 1020 0.2 2.3 0.0 106 19 14.8 35.6 39.7 13.2 19.0 0.12 0.69 6.2 12.4 45.8 35.6 14.2 4.3
445 1030 0.2 2.0 0.0 92 22 14.8 37.0 37.3 13.3 18.4 0.12 0.69 6.3 12.1 44.7 37.0 13.5 4.1
445 1040 0.2 1.8 0.0 108 19 14.7 35.8 39.3 13.3 18.4 0.22 0.79 6.4 12.5 45.4 35.8 13.4 4.3
445 1030 0.2 1.9 0.0 127 22 14.7 37.6 36.3 13.3 18.4 0.13 0.70 7.3 9.0 46.1 37.6 13.6 4.2
CBU-14 435 1030 0.2 2.3 0.0 246 27 13.8 43.3 26.6 12.9 20.9 0.82 0.97 3.9 8.0 44.7 43.3 13.7 4.5
435 1020 0.2 3.1 0.0 251 27 13.6 40.1 32.0 13.2 19.0 0.29 0.44 3.8 10.9 45.3 40.1 13.7 4.2
435 1010 0.7 2.7 0.0 328 26 13.5 43.1 27.0 13.2 .19.0 0.07 0.22 3.8 8.6 44.5 43.1 13.5 4.3
435 1010 0.7 2.7 0.0 291 30 13.5 41.1 30.3 13.5 17.2 0.07 0.22 4.1 11.2 43.6 41.1 13.4 4.3
CBU-15 425 1030 0.7 2.8 0.0 392 35 13.3 43.8 25.8 13.1 19.6 0.13 0.17 3.5 7.4 45.3 43.8 13.2 4.3 ~
425 1040 0.7 2.6 0.0 351 34 13.5 43.8 25.8 13.4 17.8 0.14 0.18 3.5 9.3 43.4 43.8 13.2 4.3 -...I
425 1040 0.7 2.9 0.0 388 35 13.3 42.0 28.8 13.4 17.4 0.14 0.18 3.1 10.0 44.9 42.0 13.6 4.3 -...I
00
425 1070 0.7 3.0 0.0 317 27 13.6 41.5 29.7 13.4 17.8 0.02 0.06 3.8 8.4 46.3 41.5 13.9 4.3 VI
CBU-16 415 1020 0.7 3.9 0.0 714 40 13.2 47.0 20.3 12.9 20.9 0.06 NO 4.7 6.9 41.8 47.0 13.1 4.4 00
425 1030 0.7 3.4 0.0 319 25 13.6 41.9 29.0 13.3 18.4 0.19 NO 4.3 8.9 44.9 41.9 13.2 4.4 0\
CBU-17 435 1020 0.7 3.7 0.0 333 29 13.6 43.7 26.0 13.7 16.0 0.10 NO 2.7 10.6 43.1 43.7 13.1 4.0
435 2010 0.7 8.8 0.0 73 12 15.3 28.1 52.4 10.7 34.4 0.04 NO 3.5 23.2 45.2 28.1 12.2 4.1
445 1040 0.7 4.5 0.0 224 26 13.6 36.2 38.6 14.0 14.1 0.17 NO 3.3 19.2 41.3 36.2 14.0 4.3
445 2020 0.7 11.5 0.0 41 9 14.4 24.2 58.9 9.6 41.1 0.02 NO 2.9 27.4 45.5 24.2 11.4 4.0
445 1980 0.7 3.4 0.0 39 15 15.9 28.4 49.2 10.5 35.6 0.01 NO 9.0 15.2 47.3 28.4 12.4 3.9
CBU-18 415 2010 0.7 9.8 0.0 664 52 13.0 45.8 22.4 13.1 19.6 0.05 NO 2.9 7.5 43.7 45.8 13.4 4.2
415 2480 0.7 11.3 0.0 484 44 13.3 45.5 22.8 13.2 19.0 0.06 NO 3.2 6.5 44.7 45.5 13.5 4.3
415 2520 0.7 6.9 0.0 928 64 12.9 48.2 18.2 12.9 20.9 0.05 NO 3.2 5.6 43.0 48.2 12.9 4.2
425 2000 0.7 6.7 0.0 259 29 13.5 42.3 28.2 13.6 16.6 0.05 NO 3.3 7.1 47.2 42.3 13.2 4.2
CBU-19 1.8 50,300 741
Feed 415 970 1.8 4.0 0.7 4,370 153 12.9 53.2 9.8 13.1 19.6 0.02 NO 1.5 3.1 42.2 53.2 12.7 4.5
415 1960 1.8 4.5 1.4 1,510 87 14.4 52.5 11.0 12.3 24.5 0.00 NO 2.8 3.7 41.1 52.5 12.5 4.3 N
425 1030 1.8 3.0 0.7 1,420 82 13.3 52.7 10.7 12.6 22.7 0.00 NO 1.6 4.6 41.2 52.7 13.1 4.3
N
425 2030 1.8 4.2 0.8 606 45 12.7 47.2 20.0 12.3 24.5 0.07 NO 3.1 6.1 43.6 47.2 13.2 4.3
435 1060 1.8 1.8 1.1 615 49 14.2 47.9 18.2 12.5 23.3 0.07 NO 3.5 5.0 43.6 47.9 13.0 4.3
435 2000 1.8 4.3 0.3 269 37 13.3 41.1 30.5 12.6 22.7 0.22 NO 8.5 4.7 45.6 41.1 13.6 4.0
CBU-20 1.5 46,000 737
Feed 445 2040 1.5 3.4 0.0 72 14 15.0 32.2 45.5 10.8 33.7 0.01 NO 7.7 13.5 46.7 32.2 12.0 3.9
445 1050 1.5 2.1 0.1 422 40 13.5 46.1 21.9 12.8 21.5 0.35 NO 4.7 6.2 43.0 46.1 13.7 4.3
445 2040 1.5 3.1 0.0 94 16 15.5 36.2 38.7 11.7 28.2 0.27 NO 7.1 9.2 47.6 36.2 12.8 4.2
445 2030 1.5 2.3 0.0 345 32 13.6 43.2 26.8 12.9 20.9 0.15 ND 4.3 5.9 46.7 43.2 13.8 4.4
CUU-21 10.8
Feed 435 2000 10.8 2.4 6.2 2,170 105 10.7 51.0 13.6 12.3 24.5 0.07 NO 3.3 4.2 41.5 51.0 12.5 4.5
435 1030 10.8 1.4 6.7 2,110 145 13.5 50.0 15.3 13.0 20.2 0.11 NO 3.6 4.0 42.4 50.0 13.0 4.1
CllU-23 10.8 56,200 763
TABLE Ie-continued
COLD LAKE HEAVY OILS RUN DATA
Feed 445 2010 10.8 2.3 2.8 228 33 15.1 34.2 42.0 13.1 19.6 0.31 0.93 9.7 7.7 48.4 34.2 13.1 3.7
445 2020 10.8 2.4 7.1 202 29 14.5 37.1 37.1 12.6 22.7 0.45 1.07 8.2 8.3 46.5 37.1 14.1 3.7
445 2010 10.8 3.8 2.0 196 36 15.6 35.5 39.8 12.4 23.9 0.12 3.73 5.5 10.4 48.6 35.5 13.1 4.0
445 2000 10.8 2.9 1.7 225 33 14.6 36.7 37.8 11.6 28.8 0.12 2.68 6.2 10.5 46.5 36.7 13.2 4.0
445 1970 10.8 3.9 0.4 242 20 14.7 38.3 35.1 13.2 19.0 0.21 0.83 3.5 13.9 44.4 38.3 13.6 3.7
CBU-24 9.7 58,700 751
Feed 435 2040 9.7 2.8 6.4 748 70 13.8 44.9 23.9 12.2 25.2 0.02 0.75 4.6 5.3 45.1 44.9 13.0 3.8
435 2020 9.7 2.6 9.0 688 78 13.2 44.2 25.1 11.9 27.0 0.01 0.74 3.2 9.9 42.7 44.2 12.4 3.8
435 2070 9.7 2.5 7.1 740 80 12.2 49.0 16.9 13.0 20.2 0.10 0.83 4.2 4.1 42.7 49.0 13.1 3.9
435 2000 9.7 2.8 5.9 756 79 12.7 47.5 19.5 11.7 28.2 0.00 0.73 3.8 3.3 45.4 47.5 13.5 3.7 ~
CBU-25 9.8 66,500 818 ~
Feed 425 2010 9.8 3.5 4.3 1,030 80 12.9 47.2 20.0 12.3 24.5 0.08 0.04 4.0 4.1 44.8 47.2 12.8 4.1
425 2030 9.8 2.8 7.8 1,110 81 12.7 50.3 14.7 13.2 19.0 0.08 0.08 2.5 5.3 41.9 50.3 12.8 4.1
425 2050 9.8 3.0 4.3 1,040 79 12.7 51.9 12.0 12.4 23.9 0.11 0.09 3.3 2.4 42.5 51.9 13.1 3.9
425 2050 9.8 2.8 8.7 1,160 87 12.3 54.5 7.7 13.0 20.2 0.09 0.07 1.8 5.7 38.1 54.5 12.9 4.1
*Water- and solids-free basis.
··Viscosity measured 011 oil after coke was removed.
···Residcnce time for continuous unit was calculated for temperatures within 50 C. of reaction temperature.
Volume % Sulfur Distribution
IBP-450" F. 450- 650- 450-950" F. % % % Gas Analysis, %
Run Vol % "API Sp gr 650" F. 950" F. "API Sp gr Liquid Gas Solids H2 CH4 CO CO2 C2H6 H2S C3HS C2H4 C3H6 Other ..f:>.
Cold Lake Crude - Barrel 1 ~
-...l
Feed 5.3 31.9 .866 20.4 16.7 19.8 .935 .?O
Run I 2.3 33.2 .859 21.6 20.1 20.3 .932 Ul
Run 2 5.4 33.3 .859 20.7 18.3 19.8 .935 00
Cold Lake Crude - Barrel 2 0\
Feed 4.5 32.7 .862 21.7 17.4 19.8 .935
Run I 7.3 31.5 .868 20.1 18.9 19.4 .938
Run 2 13.5 41.2 .819 21.9 26.9 20.2 .933
Run 3 9.7 39.2 .829 22.3 24.5 19.8 .935
Continuous Unit Runs
CBU-I" 4.6 33.0 .860 18.5 22.1 20.3 .932
6.7 33.2 .859 22.1 21.4 20.0 .934
6.0 32.5 .863 19.8 22.9 19.8 .935
CBU-2"" 7.9 32.5 .863 19.5 21.4 19.7 .936
11.0 31.9 .866 20.3 21.6 19.4 .938
10.4 35.2 .849 22.0 25.5 19.8 .935 ~
11.9 40.6 .822 17.8 25.1 20.0 .934 92 9 0 Trace 33.3 0.3 7.2 20.8 22.2 16.2 ..f:>.
CBU-3 11.2 42.0 .816 21.4 24.7 20.0 .934 88 5 0 Trace 39.1 0.6 7.0 23.8 12.2 17.4
14.3 43.5 .809 24.0 25.2 19.5 .937
16.3 46.6 .794 23.6 25.7 19.8 .935 84 19 0 Trace 35.7 0.6 4.6 22.1 20.9 16.3
22.7 45.6 .799 27.6 22.2 17.8 .948 79 23 0 Trace 35.0 Trace 3.9 23.9 19.0 18.2
CBU-4 9.9 42.7 .812 19.1 26.8 21.3 .926 85 7 0 Trace 40.2 Trace 5.2 23.9 13.2 17.5
11.6 41.7 .817 23.6 25.9 22.0 .922 82 18 0 Trace 34.7 Trace 5.3 22.4 19.9 17.7
18.6 48.3 .787 27.0 24.6 20.2 .933 76 26 0 0.0 36.1 Trace 4.6 23.3 18.8 17.4
21.1 48.3 .787 26.2 25.7 19.7 .936 74 29 0 0.0 38.3 0.0 3.4 24.5 15.7 18.0
CBU-5 14.3 42.8 .812 22.6 29.0 19.7 .936 84 19 0 0.0 25.5 0.0 2.2 28.7 22.2 21.3
CBU-6 7.6 40.4 .823 20.8 26.2 21.1 .927 90 6 0 2.3 37.5 0.0 3.3 18.9 24.5 13.4
7.4 38.6 .832 20.1 23.2 20.3 .932 92 4 0 2.4 37.7 0.0 3.1 19.4 23.1 14.3
7.6 42.6 .813 20.7 27.3 21.6 .924 87 8 0 Trace 40.1 0.0 2.5 21.3 21.5 14.5
TABLE Ie-continued
COLD LAKE HEAVY OILS RUN DATA
7.0 41.8 .816 20.6 26.7 2\.1 .927 92 9 0 1.9 29.7 0.0 3.0 25.2 23.6 16.6
CBU-8 18.8 39.1 .829 24.0 21.8 21.6 .924 88 17 0 0.0 23.2 0.0 2.9 30.5 20.3 23.1
15.5 36.9 .840 22.7 21.7 19.0 .940 92 8 I 0 0.0 27.7 0.0 3.3 25.1 26.9 17.0
22.0 40.8 .821 24.1 20.7 18.1 .946 84 20 0 0.0 14.9 0.0 2.9 34.3 26.2 21.7
21.3 39.6 .827 24.3 17.7 18.7 .942 88 17 0 0.0 22.2 0.0 1.9 29.3 23.2 22.8
29.8 41.0 .820 25.9 17.8 16.5 .956 83 17 0 0.0 25.8 0.0 1.7 31.3 16.5 23.7
25.2 36.6 .842 21.6 18.3 17.5 .950 88 20 0 0.0 27.0 0.0 2.1 28.3 20.9 21.9
CBU-9 5.4 35.2 .849 20.9 22.9 21.0 .928 90 9 0 6.7 24.8 0.6 4.0 17.6 23.0 11.9 7.8 3.4
3.0 35.2 .849 23.6 19.7 21.6 .924 90 8 0 4.7 27.8 0.6 5.4 20.0 21.0 13.9 4.1 2.5
5.5 35.6 .847 16.2 26.4 22.0 .922 90 9 0 6.5 22.5 0.6 3.5 18.4 22.8 13.2 8.4 4.0 N
5.9 39.6 .827 21.2 24.7 20.7 .930 88 7 0 3.4 26.2 0.4 4.3 22.3 19.5 17.1 4.0 2.8
01
CBU-IO 7.0 39.1 .830 21.7 25.0 2\.1 .927 85 14 0 4.7 27.8 Trace 2.8 20.0 20.9 14.9 8.1 4.3
7.9 40.6 .822 22.8 25.8 20.8 .929 90 10 0 2.3 27.8 Trace 2.7 22.6 21.0 17.2 3.4 3.0
10.7 43.2 .810 22.5 26.6 21.5 .925 88 13 0 4.0 20.7 Trace 2.5 23.1 19.8 17.4 7.9 4.6
14.5 44.1 .806 25.5 25.5 20.7 .930 79 18 0 1.9 30.0 Trace 2.9 25.7 16.4 19.9 0.6 2.6
CBU·11 6.2 36.5 .842 21.0 23.4 19.5 .937 90 5 0 3.3 27.7 0.0 4.9 20.6 15.4 16.2 8.7 3.3
9.2 36.6 .842 22.8 27.5 19.2 .939 86 9 0 4.2 26.3 0.0 5.4 19.7 22.3 14.6 5.9 1.6
13.7 38.4 .833 23.3 22.9 19.2 .939 85 10 0 5.0 27.0 0.0 4.0 18.3 17.9 16.9 5.5 4.6
9.2 41.2 .819 23.0 27.3 20.2 .933 81 17 0 1.5 14.2 0.0 3.8 26.1 20.3 20.6 7.0 6.3
CBU-12 9.9 42.7 .812 22.6 25.9 20.7 .930 91 15 0 2.3 28.9 Trace 3.9 21.4 21.0 15.5 4.1 2.8
16.4 43.7 .807 25.0 24.0 19.7 .936 86 17 0 \.I 26.2 Trace 2.5 26.5 20.9 19.7 1.4 2.2
22.7 43.8 .807 27.1 26.0 29.7 .942 83 18 0 Trace 30.9 0.0 2.0 26.9 17.7 21.4 0.0 0.0 ~.f;o.
25.2 41.9 .816 22.3 24.8 17.1 .952 78 31 0 \.I 26.2 Trace 1.9 27.3 15.8 23.4 2.4 1.9 -J
CBU-13 15.1 39.9 .826 24.4 24.8 19.2 .939 90 18 1.02 0.2 31.4 0.4 1.6 25.8 19.5 20.6 0.5 0.0 -J
00
15.0 43.2 .810 23.1 25.1 20.0 .934 86 17 1.02 1.8 30.1 0.3 1.7 24.8 17.1 19.7 2.7 0.8 ~
15.5 32.3 .864 24.7 24.3 19.7 .936 90 17 1.17 2.1 30.9 0.3 1.7 24.6 16.9 19.3 3.1 2.5
UI
00
11.2 41.8 .817 24.7 25.6 ,20.7 .930 86 18 1.03 2.1 30.1 0.2 1.7 24.4 16.9 23.5 2.7 0.0 0\
CBU-14 9.7 39.4 .828 20.9 26.7 :20.5 .931 96 11 2.58 1.0 31.3 Trace 1.8 24.3 20.3 18.4 2.5 0.4
13.2 40.3 .823 23.4 24.9 19.5 .937 91 12 1.17 1.9 30.1 0.3 1.9 23.6 20.0 18.0 1.6 2.7
10.5 40.8 .821 22.4 25.4 20.3 .932 91 11 0.59 0.7 30.8 0.3 2.2 23.7 19.5 18.3 1.7 2.8
13.5 38.0 .835 23.0 23.5 19.2 .939 91 12 0.56 1.0 31.2 0.3 2.2 24.0 19.7 19.2 1.7 0.7
CBU-15 8.9 40.0 .825 23.3 25.1 20.3 .932 92 10 0.38 0.0 29.2 0.4 2.7 24.4 22.5 18.2 0.0 0.0
12.0 39.4 .828 22.9 24.9 19.4 .938 92 10 0.39 0.0 26.8 0.2 2.6 27.6 21.7 19.3 0.0 0.0
11.2 38.5 .833 21.6 24.6 20.2 .933 92 9 0.40 1.9 31.0 0.4 2.5 24.0 20.5 17.4 0.0 0.0
10.1 39.4 .828 24.3 25.1 20.0 .934 92 11 0.13 2.0 30.5 0.3 2.4 23.4 21.4 17.5 0.0 0.0
CBU-16 8.3 38.7 .832 21.8 22.9 20.5 .931 94 9 0 2.3 26.1 0.8 4.6 22.3 19.8 18.5 2.9 2.5
10.9 41.4 .818 21.4 26.6 20.5 .931 94 10 0 Trace 33.3 Trace 2.4 23.9 20.6 17.4 1.8 0.0
CBU-17 13.0 42.0 .816 21.9 24.2 20.5 .931 85 11 0 Trace 32.2 0.3 2.5 25.0 21.4 18.7 0.0 0.0
27.5 38.3 .833 27.1 19.6 16.8 .954 87 8 0 Trace 34.5 0.3 1.7 27.3 15.1 21.2 0.0 0.0 N
22.9 37.8 .836 22.7 20.6 17.3 .951 92 8 0 1.6 33.0 0.3 2.2 26.4 15.8 20.9 0.0 0.0
Q'\
32.5 37.8 .836 28.8 17.9 15.0 .966 85 6 0 Trace 34.1 Trace 1.5 29.5 11.3 23.7 0.0 0.0
18.9 40.4 .823 28.0 22.9 16.8 .954 79 22 0 Trace 34.5 0.3 1.5 27.9 14.0 21.8 0.0 0.0
CBU·18 9.0 36.4 .843 23.8 23.1 20.0 .934 89 II 0 Trace 37.1 0.1 2.6 23.4 20.2 16.7 0.0 0.0
Feed 7.8 38.1 .834 21.7 26.3 20.5 .931 91 10 0 Trace 34.1 0.1 3.2 24.9 19.8 17.9 0.0 0.0
6.7 37.6 .837 21.6 24.7 20.5 .931 89 12 0 Trace 34.3 Trace 2.9 23.4 22.5 16.9 0.0 0.0
8.5 38.1 .835 23.4 27.1 20.3 .932 89 10 0 Trace 33.9 0.1 2.7 25.8 19.0 18.4 0.0 0.0
CBU-19 3.6 35.9 .845 23.9 20.9 20.7 .930 97 5 0 2.9 30.9 0.8 4.8 20.4 21.4 14.8 3.7 0.0
Feed 4.2 37.3 .838 19.0 23.6 21.0 .928 92 7 0 3.5 31.9 0.8 5.0 19.8 21.9 12.9 4.4 0.0
5.3 37.1 .839 18.9 24.1 20.8 .929 92 7 0 2.8 30.9 0.4 3.9 21.4 22.1 14.8 3.7 0.0
7.1 38.5 .832 22.4 23.4 20.8 .929 92 6 0 1.1 27.7 0.1 4.5 24.7 24.1 16.9 0.9 0.0
5.9 39.4 .828 21.1 25.2 21.4 .925 92 7 0 4.2 29.9 0.5 3.4 21.7 20.8 15.0 4.6 0.0
5.8 39.0 .830 23.2 26.4 21.3 .926 83 15 0 2.1 32.0 0.2 3.2 24.0 20.2 18.7 1.8 0.0
TABLE Ie-continued
COLD LAKE HEAVY OILS RUN DATA
CBU-20 16.5 39.9 .826 25.8 24.7 19.2 .939 80 22 0 3.3 29.6 0.0 2.4 27.4 16.3 2\.1 Trace 0.0
Feed 7.2 37.8 .836 20.7 24.5 2\.0 .928 92 10 0 \.I 31.2 0.3 3.1 24.3 2\.1 18.8 Trace 0.0
I\.2 40.4 .823 26.7 24.5 20.5 .931 88 17 0 \.I 32.9 0.0 2.6 25.9 17.4 20.2 Trace 0.0
7.0 39.4 .828 23.1 26.1 20.5 .931 93 12 0 3.2 30.3 0.4 2.9 23.8 2\.5 17.8 Trace 0.0
CBU-21 5.1 4\.5 .818 20.8 23.8 20.5 .931 97 3 0 6.7 3\.7 0.9 3.0 22.7 19.3 15.7 0.0 0.0
Feed 4.8 40.0 .825 20.6 24.7 20.3 .932 87 12 0 8.4 30.9 \.6 2.7 20.2 22.0 13.7 0.0 0.0
CBU-23 9.8 43.0 .810 26.1 27.4 20.5 .931 75 29 0 5.0 3\.8 \.I 2.8 15.8 19.8 11.7 0.0 0.0 11.9
Feed 10.2 4\.6 .817 25.1 25.2 20.3 .932 82 18 0 5.8 34.5 0.8 3.6 18.8 16.9 13.9 0.0 0.0 5.6
12.5 42.1 .815 24.5 26.2 20.0 .934 84 15 0 5.0 30.8 0.6 3.2 16.2 19.6 11.8 0.0 0.0 14.2
12.6 39.9 .825 25.3 23.9 19.7 .936 84 15 0 4.9 31.5 0.6 3.3 15.9 19.4 1\.7 0.0 0.0 12.6 ~
16.7 43.9 .807 22.6 23.8 20.3 .932 78 16 0 5.6 30.6 0.6 3.9 17.3 22.1 13.0 0.0 0.0 6.9 ~
CBU-24 6.3 38.3 .833 18.6 29.1 20.0 .934 82 II 0 4.1 30.8 2.5 3.1 13.8 23.9 9.5 0.0 1.0 12.1
Feed I\.9 40.8 .821 2\.7 23.2 19.7 .936 81 15 0 5.4 30.0 \.I 3.9 14.1 26.1 10.0 0.0 0.7 8.6
4.9 40.0 .825 2\.6 23.9 20.8 .929 82 13 0 5.6 29.7 \.5 4.1 14.0 25.5 9.6 0.0 0.7 9.3
4.0 40.3 .824 24.2 24.9 20.8 .929 79 14 0 5.2 29.4 1.3 4.1 13.9 25.2 9.8 0.0 0.8 10.4
CBU-25 4.9 38.2 .834 19.4 28.7 20.7 .930 87 12 0 4.5 27.3 1.7 4.4 14.0 27.3 9.4 0.0 0.8 10.8
Feed 6.3 39.7 .826 19.7 25.2 2\.0 .928 88 8 0 4.5 27.3 1.7 4.4 14.0 27.3 9.4 0.0 0.8 10.8
2.8 39.9 .826 18.9 26.7 2\.5 .925 88 7 0 5.4 30.4 1.4 4.9 14.1 26.0 9.2 0.0 0.7 7.9
6.7 36.6 .842 15.0 26.0 2\.1 .927 88 8 0 5.4 30.4 1.4 4.9 14.1 26.0 9.2 0.0 0.7 7.9
·Samples 2, 3 and 4
"Samples 2, 4. 6 and 8
~.f;>.
-..J
-...I
.?O
VI
00
0\
~
QC
29
4,778,586
30
Structural analysis for the Cold Lake feed and the TABLE IF-continued
CBU-6 product is given in Table ID.
An analysis was performed on the combined product CBU-15, IBP-285' F. MASS SPECTROMETER ANALYSIS
of the four CBU-15 runs. The results are given in Table C-Number Mol % Wt% Vol %
IE. Results from mass spectrometer analysis of the 5 10 .81 1.19 1.12
IBP-285° F. and 2850 F.-430' F. fractions of the CBU- 11 .15 .24 .22
Sum 68.53 69.38 71.25 15 run are given in Tables IF and IG, respectively. Olefins
TABLE 10 4 .36 .21 .21
5 4.13 2.99 3.04
STRUCTURAL ANALYSIS OF COLD LAKE CRUDE 10 6 7.30 6.34 6.37
OIL AND COLD LAKE CRUDE PRODUCTS 7 2.45 2.49 2.47
FROM CONTINUOUS-FLOW UNIT RUN CBU-6 8 1.13 1.31 1.28
Crude CBU-6 Sum 15.36 13.32 13.37
Oil Run-I Run-2 Run-3 Run-4 Cyclic Olefins
Run temperature, 'C. 415 415 425 425 6 .60 .51 .44
Residence time, min 4.1 2.7 4.3 2.7 15 7 .49 .49 .42
Structure: 8 .24 .27 .24
Light fractions Sum 1.33 1.27 1.10
I-Ring Napthenes
Paraffins 10.6 14.6 15.7 16.1 13.7
Cycloparaffins 8.9 14.7 14.6 15.2 14.6 6 2.63 2.28 2.09
Condensed cyclo- 27.6 26.0 25.8 24.3 22.8 7 4.71 4.77 4.31
paraffins 20 8 4.03 4.66 4.17
Alkyl benzenes 6.0 7.0 7.9 7.3 9.8 9 1.39 1.81 1.60
Benzo cyclo- 5.3 4.9 4.7 4.3 4.2 10 .60 .86 .76
paraffins 11 .13 .21 .19
Benzo dicyclo- ---H... --.£. ---±2.- ---±:Q.. Sum 13.49 14.60 13.12 -±:Q.. Alkyl Benzenes
paraffins 63.8 70.7 72.6 71.2 69.1
Aromatic Fractions 25 6 .06 .05 .04
7 .10 .09 .08
2-ring aromatics 13.7 10.2 11.1 11.0 11.3 8 .81 .88 .71
3-ring aromatics 5.8 4.8 4.2 4.5 5.7 9 .33 .41 .33
4-ring aromatics 0.6 2.8 1.8 3.1 3.3 Sum 1.29 1.43 1.16
5-ring aromatics 0.3 1.7 1.3 2.1 2.3
Polyaromatics 0.1 0.8 0.4 0.4 0.5 Uncorrected Specific Gravity, 20' C. = .7043
Sulfur aromatics -1:!. ---±:Q.. -.U... --l:§.. ..2Q.. 30 Specific Gravity, Corrected for S, IS' C. = 0.726
Specific Gravity, Observed, IS' C. = 0.7351
29.9 24.3 21.9 24.7 26.1
Remainder .....&l... ~ -21... --!.L ~
100.0 100.0 100.0 100.0 100.0
TABLE IE
ANALYSES ON CBU-15 COMBINED PRODUCT, RUNS 1-4
1.65
0.78
-100
42.1
14.6
Temp. Range, 'F. at 1 Atmos. Whole Oil IPB-285 285-430 430-525
Cut Vol % of Whole Oil 100 1.18 6.00 9.40
~ Vol % OH at Cut End 100 1.18 7.18 16.58
Cut Wt % of Whole Oil 100 0.89 4.86 8.19
~ Wt % OH at Cut End 100 0.89 5.75 13.94
'API Gravity 60/60 13.2 61.0 47.1 34.7
Specific Gravity 60/60 0.9782 0.7351 0.7921 0.8514
Sulfur, wt % 4.02 1.66 2.36 2.40
Nitrogen, wt %
Pour Point, 'F.
Cetane Index(2)
Smoke Point, mm
Con Carbon Res, wt %
Viscosity,
100' F., cst
210' F., cst
275' F., cst
Nickel, wppm
Vandium, wppm 162
525-650 650-950
15.52 35.03
32.10 67.13
14.32 34.66
28.26 62.92
25.2 14.7
0.9028 0.9679
2.57 3.59
297 ppm 0.22
-75 40
39.2 23.4
<10 (3)
0.63 >
4.34 99.6
1.44 7.63
8.0
ND
950+
32.87(1)
100.00(1)
37.08(1)
100.00(1)
-3.1
1.1016
5.62
37.5
10,400
192
408
*May include cyclic olefins and certain sulfur compounds.
ND None detected.
MASS SPECTROMETER ANALYSIS OF
285-430' F. FRACTION OF THE CBU-15 RUN
Sulfur balance closure = 100.1%; Vanadium closure = 93.4%.
ND = None Detected.
(l)By difference to give 100% recovery since loss is primarily in the residue.
(2}Ca1culated from midpoint of distillation fractions. not from a separate D-86 distillation.
(31Material would not wick, test not applicable.
_______T_A_BL_E _IF_______ 60
CBU-15, IBP-285' F. MASS SPECTROMETER ANALYSIS
C-Number Mol % Wt % Vol %
Paraffins
4 .89 .53 .66
5 10.98 8.17 9.26 65
6 15.19 13.50 14.40
7 19.43 20.09 20.48
8 15.46 18.22 17.97
9 5.62 7.43 7.14
TABLEIG
Paraffins
Olefins
Cycloparaffins
Condo Cycloparaffins*
Alkyl Benzenes
47.9 vol %
ND
35.3
12.7
~
100.0 vol %
4,778,586
TABLE 2B-continued
32
BOSCAN CRUDE, IBP·285' F.
MASS SPECTROMETER ANALYSIS
C·Number Mol % Wt %
.11
.43
31.26
Vol %
.13
.52
33.75
.17
.60
Sum 32.60
Alkyl Benzenes
6
7
31
EXAMPLE 2
Continuous-flow unit experiments were conducted
on the Boscan crude oil sample. An analysis of the feed
for each of these runs is given in Table 2A. Results from 5
mass spectrometer analysis of the IBP-285° F. and 285°
F.-430° F. fractions of the feed for these runs is given in
Tables 2B and 2C, respectively.
TABLE2A
Temp. Range, 'F. at I Atmos.
ANALYSES ON BOSCAN CRUDE
Whole IBP- 285- 430-
Oil 285 430 525
525
650
650950
950+
4.99 68.2
1.58 6.75
5,580
11.0 164
ND 1,216
Cut Vol % of Whole Oil
~ Vol. % OH at Cut End
Cut Wt % of Whole Oil
~ Wt % OH at Cut End
,API Gravity 60/60
Specific Gravity 60/60
Sulfur, wt %
Nitrogen, wt %
Pour Point, 'F.
Cetane Index(2)
Smoke Point, mm
Con Carbon Res, wt %
Viscosity,
100' F., cst
210' F., cst
275' F., cp
Nickel, wppm
Vanadium, wppm
100 2.29 3.29 2.59
100 2.29 5.58 8.17
100 I.73 2.62 2.24
100 1.73 4.35 6.59
11.3 58.7 47.4 33.2
0.9907 0.7440 0.7911 0.8589
5.21 0.37 1.27 3.02
-50
39.7
(3)
2.64
1.09
6.96
15.13
6.26
12.85
27.5
0.8901
3.89
239 ppm
o
42.4
12.0
27.44
42.57
26.11
38.96
18.6
0.9424
4.54
0.16
80
27.7
(3)
0.33
57.43(1)
100.00(1)
6\.04(1)
100.00(1)
2.4
1.0566
6.06
27.6
Sulfur balance closure = 100.5%.
ND = None Detected.
(I)By difference to give 100% recovery since loss is primarily in the residue.
(2)Calculated from midpoint of distillation fractions, not from a separate 0-86 distillation.
(3)Material would not wick, test not applicable.
8 1.37 1.38 1.15
9 .29 .33 .28
35 Sum 2.44 2.36 1.97
Uncorrected Specific Gravity, 20' C. = .7288
Specific Gravity, Corrected for S, IS' C. = 0.7390
Specific Gravity, Observed, IS' C. = 0.7441
TABLE2B 40 TABLE 2C
BOSCAN CRUDE, IBP-285' F. MASS SPECTROMETER ANALYSIS OF
MASS SPECTROMETER ANALYSIS 285-430' F. FRACTION OF THE BOSCAN FEED
C-Number Mol % Wt% Vol % Paraffins 60.6 vol %
Paraffins Olefins ND
5 5.21 3.54 4.15 45 Cycloparaffins 32.5
6 15.44 12.56 13.84 Condo Cycloparaffins 2.8
7 17.13 16.20 17.08 Alkyl Benzenes ~
8 14.61 15.75 16.06 100.0 vol %
9 8.26 9.99 9.93 NO None detected.
10 3.98 5.35 5.21
II .34 .50 .48 50
Sum 64.96 63.89 66.77 An analysis of the products is given in Table 2D.
l·Ring Napthenes Batch autoclave runs were also conducted on Boscan
6 3.85 3.06 2.89 crude oil. The results of these runs and further batch
7 11.50' 10.65 9.95 autoclave runs are given in Table 2E. Also, the struc-
8 7.48 7.92 7.33 tural analysis of a continuous-flow unit run of the Bos- 9 6.43 7.66 7.03 55
10 3.18 4.21 3.83 can heavy oil was determined. The results were pres-
II .17 .25 .23 ented in Table 2F.
60
65
TABLE2D
BOSCAN HEAVY OILS RUN DATA
Pres- Feed Product Solid Coke Gas IBP- 450- Resid Con- Sulfur
Temp. sure, H2O Time H2O Viscosity·· Gravity Residual Asphaltene* WI. WI. WI. 450· F. 950· F. +950 F.. Carbon WI.
Run .c. psig % min··· % cp 25· C. cp 80· C. ·API WI. % ,Conv. % Wt.% Alter. % % % % Wt.% Wt.% WI. % Wt.% %*
(Barrel I) - Batch Runs
Feed 0.9 59,300 827 11.4 68,8 20.1 0.2 4.4 26.6 68.8 14.3 5.6
Continuous Unit Runs (Barrell)
CBU·26 400 1000 0.9 3.1 0.5 3,890 161 12.3 60.6 11.9 17.1 14.9 0.02 NO 4.2 6.1 29.1 60.6 14.1 5.1
400 2020 0.9 5.2 0.6 3,150 133 13.2 56.0 18.6 17.1 14.9 0.0\ NO 2.2 7.8 34.1 56.0 14.2 5.0
415 2040 0.9 3.2 0.0 823 54 13.2 49.0 28.8 17.2 14.4 0.11 NO 4.5 8.4 38.1 49.0 15.0 4.8 W
415 1040 0.9 2.3 0.0 845 61 14.7 50.2 27.0 17.5 12.9 0.07 NO 4.5 9.8 35.5 50.2 15.0 4.9 W
425 1080 0.9 2.5 0.3 522 40 14.2 43.2 37.3 16.7 16.9 0.34 NO 4.3 15.2 37.4 43.2 15.0 4.9
CBU-27 425 2040 0.9 2.9 0.0 712 40 14.7 46.0 33.1 11.7 41.8 0.24 NO 5.2 9.8 39.0 46.0 15.5 4.6
435 2010 0.9 2.7 0.0 56 16 17.4 34.6 49.8 11.9 40.8 0.03 NO 7.0 16.4 42.0 34.6 13.3 4.8
435 1060 0.9 2.2 0.0 275 40 14.8 42.8 37.8 15.3 23.8 0.10 NO 6.0 11.3 39.9 42.8 15.7 5.0
445 1050 0.9 2.0 0.0 55 17 17.6 32.9 52.2 13.3 33.8 0.27 NO 7.4 17.5 42.2 32.9 13.2 4.6
CBU·28 435 1010 0.9 1.9 489 40 14.4 16.8 16.4 0.21 2.61 4.9
435 1030 0.9 2.1 0.0 250 28 15.7 40.2 41.6 15.9 20.9 0.10 2.50 6.3 15.3 38.3 40.2 15.2 4.8
435 1030 0.9 2.3 216 20 16.0 15.8 21.4 0.12 2.52 5.1
435 1050 0.9 2.3 0.1 251 26 15.3 40.4 41.3 15.9 20.9 0.13 2.53 5.9 14.8 39.0 40.4 14.4 4.5
CBU-29 425 1060 0.9 2.6 568 53 13.9 17.2 14.4 0.14 0.20 5.2
425 HMO 0.9 2.4 0.1 622 45 13.8 45.3 34.2 17.0 15.4 0.17 0.23 4.0 12.0 38.8 45.3 15.7 4.9 ~~
425 1040 0.9 2.4 617 46 13.8 17.2 14.4 0.17 0.23 5.0 -...l
425 1040 0.9 2.6 0.0 629 51 13.8 49.3 28.3 17.3 14.0 0.04 0.10 4.8 9.3 36.7 49.3 15.7 4.9 -...l
CBU-30 415 1030 0.9 2.7 0.0 869 59 14.8 49.0 28.7 17.0 15.4 0.13 0.13 3.6 9.9 37.5 49.0 15.3 5.0 ,po
415 1010 0.9 2.5 0.0 992 62 14.8 52.3 24.0 17.3 13.9 0.12 0.12 3.4 7.1 37.2 52.3 15.0 4.8 VI
415 1000 0.9 2.6 0.0 874 56 13.6 47.4 30.7 17.8 11.4 0.13 0.13 4.4 9.2 39.0 47.4 15.3 5.3
00
0\
415 1020 0.9 2.6 0.0 898 61 13.6 52.3 24.0 17.6 12.4 0.08 0.08 3.5 8.1 36.0 52.3 15.2 5.0
CBU·31 415 1030 0.9 4.3 0.0 775 52 13.6 48.6 29.4 17.7 11.9 0.45 NO 4.5 7.0 40.0 48.6 16.2 4.8
425 1050 0.9 4.6 0.0 706 45 13.6 45.6 33.7 17.4 13.4 0.18 NO 5.0 8.7 40.7 45.6 15.5 4.6
425 540 0.9 4.5 0.0 1,120 70 13.3 53.7 22.0 18.2 9.5 0.35 NO 4.3 6.6 35.4 53.7 15.4 5.0
435 1020 0.9 6.6 0.0 642 40 13.5 45.6 33.7 17.5 12.9 0.32 NO 3.0 12.4 39.0 45.6 15.9 4.5
CBU-35 415 500 0.9 2.5 0.0 3,335 152 13.2 56.2 18.3 17.5 12.9 0.09 NO 2.6 7.1 34.1 56.2 14.4 5.3
425 540 0.9 1.6 0.0 975 60 12.2 52.2 24.1 17.2 14.4 0.11 NO 4.6 7.9 35.3 52.2 15.0 5.3
435 550 0.9 1.6 0.0 707 73 15.4 42.3 38.5 17.3 13.9 0.25 NO 5.4 13.3 39.0 42.3 15.5 4.8
435 270 0.9 0.8 0.0 978 60 12.5 50.0 27.3 17.6 12.4 0.09 NO 5.0 8.9 36.1 50.0 15.8 5.0
CBU-36 400 1060 2.5 3.2 0.7 14,700 420 12.6 63.3 8.0 17.8 11.4 0.07 NO 2.6 4.9 29.1 63.3 14.2 5.6
415 1030 2.5 3.1 0.2 4,430 177 13.0 61.6 10.5 17.1 14.9 0.04 NO 2.4 5.7 30.4 61.6 15.1 5.1
425 1060 2.5 2.0 0.5 1,260 124 14.2 49.4 28.2 17.1 14.9 0.10 NO 5.3 7.8 37.5 49.4 15.2 5.1 W
435 1020 2.5 1.9 0.0 822 109 14.2 46.7 32.1 17.3 13.9 0.19 NO 7.0 6.7 39.6 46.7 16.0 4.8 ~
IBF-450· F. Volume % 450-950·F. Sulfur Distribution Gas Analysis, %
Run Vol % ·API Sp gr 450-650· F. 650-950· F. •API Sp gr % Liquid % Gas % Solids H2 CH4 CO C02 C2H6 H2S CJHH Other
(Barrel 1) - Batch Experiments
Feed 5.5 47.3 .792 11.5 17.4 23.7 .912
Continuous Unit Runs (Barrel I)
CBU-26 7.4 42.6 .813 13.9 17.5 23.0 .916 88 18 0 4.9 32.7 0.5 5.1 13.1 26.9 5.9 10.9
9.4 43.0 .811 15.4 20.8 21.8 .923 88 7 0 3.3 27.8 0.4 4.7 13.1 30.0 7.4 13.3
10.2 43.5 .809 17.9 22.7 22.3 .920 84 13 0 1.6 22.3 0.1 3.5 13.7 30.5 8.9 19.4
11.9 41.9 .816 18.1 20.0 21.8 .923 87 10 0 1.8 23.1 0.8 3.3 12.6 27.4 8.1 22.9
17.8 37.9 .835 18.0 21.3 20.7 .930 85 14 0 2.3 27.2 0.1 3.0 13.9 28.0 9.5 16.0
CBU-27 12.3 47.3 .791 18.4 24.1 22.8 .917 80 17 0 1.6 3ll.4 Traee 3.3 16.3 29.4 11.4 7.6
TABLE 2D-continued
BOSCAN HEAVY OILS RUN DATA
20.3 43.5 .808 22.7 22.7 21.1 .927 81 23 0 3.3 25.9 Trace 2.6 15.7 25.8 11.0 15.7
14.0 42.5 .813 22.5 20.8 21.6 .924 86 17 0 1.6 31.2 Trace 2.5 17.7 24.6 13.0 9.4
21.6 43.7 .808 22.5 22.8 20.5 .931 78 20 0 2.0 30.1 0.1 2.2 17.4 22.1 12.0 14.1
CBU-28 I 2.6 27.4 0.1 2.7 14.3 26.4 10.0 16.4
18.8 42.0 .815 21.0 20.1 20.7 .930 82 17 0 3.5 22.7 Trace 2.4 15.3 24.9 10.9 20.3
1.8 29.2 Trace 2.5 18.2 26.4 9.0 12.7
18.5 44.7 .803 19.0 23.3 21.1 .927 77 19 0 1.6 30.4 0.1 2.6 16.3 25.3 11.2 12.5
CBU·29 Trace 29.5 0.2 4.8 15.5 32.1 10.6 7.2
14.7 42.1 .815 20.4 21.2 21.0 .928 85 14 0.2 1.3 29.2 0.1 3.3 15.2 31.5 10.4 8.9
1.6 30.8 0.2 3.3 15.3 32.0 10.5 5.9 CH
11.2 42.5 .813 17.4 21.4 20.8 .929 86 15 0.1 1.6 28.6 0.1 3.4 15.3 32.0 10.4 8.5 UI
CBU-30 11.9 42.0 .815 18.9 21.1 21.5 .925 88 11 0 1.3 29.4 Trace 3.1 13.4 32.2 9.6 10.9
8.8 44.6 .804 18.0 21.9 22.6 .918 85 11 0 1.5 29.0 0.1 2.5 13.4 32.3 9.9 11.2
11.1 41.3 .819 20.5 21.5 21.5 .925 92 13 0 1.5 28.9 0.1 2.9 13.4 32.6 9.8 10.8
10.0 42.9 .811 16.8 22.2 22.5 .919 87 12 0 1.5 27.9 0.2 3.0 16.1 31.4 9.4 10.5
CBU·31 8.7 44.0 .806 19.4 24.2 23.0 .916 83 14 0 1.5 27.4 Trace 2.3 14.1 30.3 11.3 13.1
10.9 44.3 .805 19.4 25.3 22.8 .917 90 18 0 0.9 24.3 Trace 3.2 14.0 32.9 11.9 12.8
8.2 43.9 .807 17.1 21.0 23.3 .914 86 14 0 1.6 28.1 0.1 2.6 14.3 27.4 11.4 14.5
14.8 40.6 .822 19.6 21.4 20.8 .929 77 19 0 3.2 32.5 0.4 2.5 12.3 29.1 9.4 10.6
CBU-35 8.4 39.0 .830 18.5 17.9 21.1 .927 94 4 0 1.6 32.9 3.2 3.9 13.0 19.6 10.1 15.7
9.8 42.7 .812 16.4 22.3 22.1 .921 92 13 0 2.1 26.9 0.1 2.5 12.9 28.6 9.8 17.1
16.1 40.9 .821 22.2 19.3 20.0 .934 83 17 0 1.9 23.1 0.6 1.8 13.5 30.2 10.8 18.1 ~
11.2 43.6 .808 17.1 22.7 21.8 .923 85 17 0 6.5 30.0 2.0 1.2 11.9 24.8 7.5 16.1 :...
CBU·36 6.0 42.0 .816 12.5 18.9 22.3 .920 98 4 0 4.7 28.1 0.9 3.7 12.1 28.3 8.2 14.0 ....
6.8 38.5 .832 16.1 16.5 22.0 .922 90 7 0 4.0 27.8 1.1 3.5 12.5 28.5 8.3 14.3 .¥J
9.4 39.9 .826 17.7 22.6 21.0 .928 88 13 0 3.6 27.2 0.7 3.0 12.8 29.4 8.9 14.4 VI
8.2 41.7 .817 18.4 24.8 22.1 .921 82 16 0 4.1 27.0 0.6 2.6 12.9 26.4 11.5 14.9 00
0\
*Watcr- and solids-free basis.
··Viscosity measured on oil afler coke was removed.
···Residence time for continuous unit was calculated for temperatures within 5° C. of reaction temperature.
CH
01
Run
Feed
BO I
BO 2
CBU-7
Run
. Feed
BO I
B02
CBU-7
37
4,778,586
38
TABLE 2E
BOSCAN HEAVY OILS RUN DATA
Pres- Feed Product Viscosity" Residual Asphaltene' Solid Coke Gas IBPTemp
sure, HzO Time HzO cp cp Gravity WI. Conv. Wt. Alter. Wt. WI. Wt. 450' F.
'C. psig % min··· % 25' C. 80' C. 'API % % % % % % % Wt.%
0.8 104,900 1,510 10.1 73.6 20.9 0.00 0.2 2.6
400 460 0.8 15 Trace 1,190 87 12.2 55.5 24.6 17.8 14.8 0.00 0.0 1,4 8,4
415 760 0.8 15 Trace 118 21 15.7 40.5 45.0 15.3 26.8 0.11 2,4 4,4 12.7
415 1060 0.8 1.8 0.6 2,300 III 14.4 52.4 28.8 17.3 17.2 0.00 0.0 1.7 11.5
425 1030 0.8 1.9 Trace 1,180 81 14.1 50.6 31.3 17.3 17.2 0.00 0.0 4.0 8.5
450- Resid Con- Sulfur Volume % Sulfur Distribution
950' F. +950 F. Carbon Wt. IBF-450' F. 450- 650- 450-950' F. % % % CI
Wt.% Wt.% Wt.% %' Vol % 'API Sp gr 650' F. 950' F. 'API Sp gr Liquid Gas Solids ppm
23.6 73.6 14.0 5.6 3.0 42.5 .813 9.9 15.4 24.3 .908 7.2
34.6 55.5 14.6 5.2 10.6 49.0 .784 15.4 21.7 22.6 .918 93 0 0
40.0 40.5 13.0 4.8 15.4 47.0 .793 20.0 21.8 22.8 .917 76 16 3
34.4 52.4 14.6 5.3 13.7 40.0 .825 17.6 18.7 20.2 .933 93 9 0
36.9 50.6 15.6 5.1 10.4 43.1 .810 18.0 21.8 22.0 .912 89 14 0
Pour Point Gas Analysis, %
Run 'c. Hz CH4 CO COZ C2H6 HZS C3HS Other
Feed 18
BO I -5
BO 2
CBU-7 -10 Trace 20.9 0.0 5.4 24.2 34.6 14.9
-4 Trace 23.1 0.0 4.2 23.5 32.2 17.0
*Water- and solids-free basis.
"Viscosity measured on oil after coke was removed.
··-Run CBU·7 was run in the continuous unit. All other runs were performed in the batch autoclave.
For 10' API oil, 10 lb. salt/lOoo bbl. is equivalent to 18 ppm Cl.
TABLE2F TABLE 2F-continued
5.2
4.6
3.8
1.5
-1l...
3\.3
-l:Q....
100.0
CBU-7
6.0
5.3
3.1
I.l
7.4
31.8
6.2
100.0
BO-I BO-2
4.3
3.9
1.7
0.8
~
25.3
......2:.L
100.0
Feed
2.5
1.2
0.3
0.3
-!!.:.!...
23.0
-1l....
100.0
STRUCTURAL ANALYSES OF BOSCAN
HEAVY CRUDE OIL FEEDS AND RUN PRODUCTS
(Wt%)
3-Ring Aromatics
4-Ring Aromatics
5-Ring Aromatics
Polyaromatics
Sulfur Aromatics
Remainder
40
35
STRUCTURAL ANALYSES OF BOSCAN
HEAVY CRUDE OIL FEEDS AND RUN PRODUCTS
(Wt %)
Feed BO-I BO-2 CBU-7
Run Temperature, 'C. 400 415 425
Residence Time, Min. 15 15 1.9
Structure
Light Fractions
Paraffins 12.7 19.7 19.8 17.3
Cycloparaffins 14.8 15.6 15.0 14.8
Condensed 28.6 20.8 14.9 15.5
Cycloparaffins
Alkyl Benzenes 4.9 5.4 5.9 7.0
Benzo Cycloparaffins 3.7 3.1 3.6 3.8
Benzo Dicycloparaffins -.!Q... ~ -.bL --lL
68.7 67.6 62.0 61.7
Heavier Fractions
2-Ring Aromatics 7.6 8.8 8.9 10.9
45 An analysis was performed on the combined product
of the four CBU-30 runs. The results are given in Table
2G. Results from mass spectrometer analysis of the
IBP-285° F. and 285° F.-430° F. fractions of the CBU30
run are given in Tables 2H and 21, respectively.
TABLE2G
Temp. Range, 'F. at I Atmos.
ANALYSES ON CBU-30 COMBINED PRODUCT, RUNS 1-4
Whole IBP- 285- 430- 525
Oil 285 430 525 650
650950
950+
38.5
38.99(1)
100.00(1)
43.63(1)
100.00(1)
-2.2
1.0947
5.73
15,220
226
84.9
7.50
9.9
31.73
61.01
31.05
56.37
16.4
0.9564
4.43
0.23
90
25.3
(3)
\.20
4.55
1.49
1.66
0.81
-50
45.2
14.0
100 2.27 6.69 7.77 12.55
100 2.27 8.96 16.73 29.28
100 1.67 5.41 6.70 11.54
100 1.67 7.08 13.78 25.32
13.3 64.9 47.6 36.6 25.9
0.9771 0.7206 0.7901 0.8420 0.8990
4.79 1.09 2.34 3.02 4.06
485 ppm
o
40.2
10.8
Cut Vol % of Whole Oil
~ Vol. % OH at Cut End
Cut Wt % of Whole Oil
~ Wt % OH at Cut End
'API Gravity 60/60
Specific Gravity 60/60
Sulfur, wt %
Nitrogen, wt %
Pour Point, 'F.
Cetane Index(Z)
Smoke Point, mm
Con Carbon Res, wt %
Viscosity,
100' F., cst
210' F., cst
275' F., cp
Nickel, wppm
39
4,778,586
TABLE 2G-continued
40
ANALYSES ON CBU-30 COMBINED PRODUCT, RUNS 1-4
Whole IBP- 285- 430- 525 650-
Temp. Range, 'F. at 1 Atmos. Oil 285 430 525 650 950
Vanadium, wppm 849 3.1
Sulfur balance closure = 97.9%; Vanadium closure = 92.1%.
ND = None Detected.
(I)By difference to give 100% recovery since loss is primarily in the residue.
(2)Calculated from midpoint of distillation fractions, not from a separate 0-86 distillation.
(3)Material would not wick, test not applicable.
950+
1,573
TABLE2H
CBU-30, IBP-285' F. MASS SPECTROMETER ANALYSIS
C-Number Mol % Wt % Vol %
TABLE 2H-continued
CBU·30, IBP-285' F. MASS SPECTROMETER ANALYSIS
C-Number Mol % Wt % Vol %
Paraffins 15 Alkyl Benzenes
4 3.10 1.91 2.35 6 .06 .05 .04
5 13.49 10.31 11.65 7 .40 .39 .32
6 18.41 16.81 17.87 8 .91 1.02 .83
7 15.11 16.04 16.30 9 .27 .35 .28
8 12.32 14.90 14.65 20 Sum 1.64 1.81 1.46
9 5.20 7.07 6.77 Uncorrected Specific Gravity, ZO' C. = .7035
10 1.07 1.61 1.51 Specific Gravity, Corrected for S. IS' C. = 0.720
11 .11 .19 .17 Specific Gravity, Observed, IS' C. = 0.7206
Sum 68.82 68.83 71.28
Olefins
4 .55 .33 .34 TABLE 21 .
5 4.70 3.49 3.55 25
MASS SPECTROMETER ANALYSIS OF
6 3.93 3.50 3.51 285-430' F. FRACTION OF THE CBU-30 RUN
7 1.01 1.05 1.04
8 .40 .48 .47 Paraffins 54.4 vol %
Sum 10.59 8.85 8.90 01efins ND
Cyclic Olefins Cycloparaffins 34.7
6 .54 .47 .41 30 Condo Cycloparaffins' 6.8
7 .50 .51 .44 Alkyl Benzenes --.!!...
8 .55 .64 .55 100.0 vol %
Sum 1.59 1.62 1.40 *May include cyclic olefms and certain sulfur compounds.
I·Ring Napthenes ND None detected.
6 3.41 3.04 2.77 7 6.75 7.02 6.32 35
8 5.70 6.78 6.05 EXAMPLE 3
9 1.22 1.63 1.44
10 .28 .41 .36 Batch autoclave and continuous flow unit runs were
Sum 17.36 18.89 16.95 conducted on the Tia Juana crude sample. The results
are given in Table 3A.
TABLE3A
TIA JUANA HEAVY OILS RUN DATA
Run
Feed
TJ I
TJ 2
TJ 3
TJ4
TJ 5
CBU-33
Run
Feed
TJ 1
TJ2
TJ3
TJ4
TJ 5
CBU-33
Pres- Feed Product Viscosity" Residual Asphaltene' Solid
Temp sure, HzO Time HzO cp cp Gravity Wt. Conv. WI. Alter. WI.
'C. psig % min*·· % 25' C. 80'C. 'API % % % % %
0.0 21,100 476 12.0 64.9 12.4 0.00
350 250 0.0 15 Trace 9,740 249 12.8 59.5 8.3 13.0 -4.8 0.00
380 250 0.0 15 0.0 10,500 331 13.9 57.2 11.9 12.4 0.0 0.07
400 360 0.0 15 Trace 1,500 79 16.9 52.2 19.0 13.4 -8.1 0.02
415 560 0.0 15 0.06 925 49 15.7 39.6 39.0 14.7 -18.6 0.39
425 650 0.0 15 0.0 477 29 19.2 39.3 39.5 13.6 -9.7 0.09
415 1030 0.0 3.5 0.0 2,570 117 16.1 51.5 20.6 13.1 -5.3 0.03
425 1020 0.0 3.6 0.0 863 103 14.8 45.5 29.9 13.5 -9.1 0.06
435 960 0.0 2.7 0.0 397 46 16.4 40.7 37.3 13.4 -8.1 0.20
Coke Gas IBP- 450- Resid Con· Sulfur Volume %
WI. WI. 450' F. 950' F. +950F Carbon WI. IBP-450' F. 450- 650- 450-950' F.
% % WI. % Wt.% Wt.% Wt.% %' Vol % 'API Sp gr 650' F. 950' F. 'API Sp gr
0.02 1.6 33.5 64.9 12.2 2.8 1.9 37.0 .840 11.7 23.9 21.5 .925
0.0 0.3 6.3 38.9 54.5 11.9 2.8 7.3 35.9 .845 16.4 24.1 18.9 .941
0.0 0.2 4.9 37.7 57.2 11.8 2.7 5.7 35.8 .846 17.2 22.1 19.8 .935
0.0 1.8 6.1 39.5 52.6 12.7 2.8 7.0 38.7 .831 17.1 23.4 21.2 .927
0.0 2.2 15.0 43.2 39.6 13.5 2.7 17.3 38.9 .831 22.5 21.8 19.2 .939
1.9 0.4 9.8 48.6 39.3 13.4 2.7 11.1 39.2 .829 21.1 26.9 20.7 .930
ND 2.2 8.4 37.9 51.5 12.7 2.8 9.9 41.6 .817 16.8 22.7 20.7 .930
ND 1.4 7.8 45.3 45.5 13.5 2.7 9.3 42.1 .815 20.2 26.8 20.3 .932
ND 4.6 9.7 45.0 40.7 14.1 2.7 11.7 42.1 .815 21.0 26.6 20.3 .932
Sulfur Distribution Pour
% % % CI Point Gas Analysis, %
Run Liquid Gas Solids ppm 'C. HZ CH.j CO COz CZH6 HzS C3Hs Other
Feed 0.49 9
4,778,586
41 42
TABLE 3A-continued
TIA JUANA HEAVY OILS RUN DATA
TJI 100 0 0 6
TJ 2 96 0 0 7
TJ3 100 0 0 -3
TJ4 96 0 0 -9
TJ5 95 0 3 -13
CBU-33 95 5 0 -10 3.4 30.4 1.3 6.9 13.4 13.6 11.9 18.8
94 8 0 -19 1.9 37.0 1.0 4.8 16.0 11.6 13.4 14.3
93 10 0 -25 1.7 34.6 0.6 5.3 16.0 9.8 14.4 17.6
·Water- and solids-free basis.
··Viscosity measured on oil after coke was removed.
···Run CBU·33 was run in the continuous unit. All other runs were performed in the batch autoclave.
For 10' API oil. 10 lbs salt/IOOO bbls is equivalent to IS ppm Cl.
TABLE3B
STRUCTURAL ANALYSES OF
TIA JUANA HEAVY CRUDE OIL FEED
(Wt%)
Structural data for the Tia Juana crude oil feed is
given in Table 3B.
Batch autoclave and continuous unit runs were conducted
on the Zuata crude oil sample. The results are
given in Table 4A.
TABLE4A
3.4
1.3
0.3
0.3
_7_.2_
22.3
_6_.9_
100.0
STRUCTURAL ANALYSES OF
TIA JUANA HEAVY CRUDE OIL FEED
(Wt%)
TABLE 3B-continued
3-Ring Aromatics
4-Ring Aromatics
5-Ring Aromatics
Po[yaromatics
Sulfur Aromatics
EXAMPLE 4
Remainder
Structure
30
25
20
9.8
11.2
16.7
28.0
5.1
4.4
~
70.8
Heavier Fractions
2-Ring Aromatics
Light Fractions
Paraffins
Cycloparaffins
Condensed
Cycloparaffins
Alkyl Benzenes
Benzo Cycloparaffins
Benzo Dicycloparaffins
Structure
ZUATA HEAVY OILS RUN DATA
Pres- Feed Product Viscosity" Residual Asphaltene' Solid
Temp sure, HzO Time HzO cp cp Gravity WI. Conv. Wt. Alter. Wt.
Run ·C. psig % min**~ % 25· C. 80· C. ·API % % % % %
Feed 9.5 193.000 1.440 9.4 64.6 18.0 0.15
ZUI 400 2200 9.5 15 1.2 2,410 104 10.7 52.4 18.9 14.7 18.3 0.04
ZU2 370 1750 9.5 15 11.8 46,200 512 9.7 61.7 4.5 14.4 20.0 0.08
ZU3 360 1850 9.5 15 2.1 9,000 196 12.9 51.7 20.0 14.2 21.1 0.07
ZU4 415 2275 9.5 15 Trace 457 38 15.7 41.3 36.1 14.8 17.8 0.32
CBU-34 415 1060 9.5 0.9 10.7 29,800 . 514 12.2 56.3 12.8 18.2 -I.! 0.17
425 1020 9.5 1.4 7.3 9,410- 234. 12.2 56.2 13.0 17.1 5.0 0.16
435 1040 9.5 2.7 0.2 2,800 103 14.1 48.7 24.6 14.4 19.9 0.19
Coke Gas IBP- 450- Resid Con- Sulfur Volume %
WI. Wt. 450· F. 950· F. +950F Carbon WI. IBP-450· F. 450- 650- 450-950· F.
Run % % WI. % Wt.% Wt.% Wt.% %' Vol % ·API Sp gr 650· F. 950· F. ·API Sp gr
Feed 0.6 0.9 33.9 64.6 11.6 3.6 1.2 43.2 .810 12.3 23.9 18.9 .941
ZU 1 0.0 1.0 5.5 41.1 52.4 12.8 3.7 6.7 41.7 .817 17.3 26.3 19.5 .937
ZU2 0.0 1.9 2.8 33.6 61.7 12.5 3.8 7.5 28.4 19.5 .937
ZU3 0.0 0.7 7.6 40.0 51.7 12.8 3.4 8.8 36.5 .842 18.9 22.4 17.6 .949
ZU4 0.9 3.8 8.8 45.2 41.3 13.6 3.4 10.4 41.5 .818 21.8 24.9 19.2 .939
CBU-34 ND 3.1 2.7 37.9 56.3 12.3 3.5 3.2 35.4 .847 15.1 24.9 18.4 .944
ND 3.4 2.7 37.7 56.2 13.7 3.5 3.2 37.3 .838 14.5 25.6 19.5 .937
NO 2.9 5.1 43.3 48.7 14.4 3.2 6.1 39.2 .829 18.5 27.1 19.0 .940
Sulfur Distribution Pour
% % % Cl Point Gas Analysis, %
Run Liquid Gas Solids ppm .c. Hz CH4 CO COz CZH6 HzS C3HS Other
Feed 14.9 24
ZU I 103 0 0
ZU2 106 0 0 13
ZU 3 94 0 0 6
ZU4 94 0 I
ZU 5 95 4 0 5 3.8 37.7 Trace 8.2 15.6 11.5 12.7 10.5
CBU-34 95 5 0 2 1.7 33.1 3.4 5.2 14.0 15.0 10.9 16.7
43
4,778,586
44
TABLE 4A-continued
ZUATA HEAVY OILS RUN DATA
86 10 0 - 7 1.6 32.9 3.2 3.9 13.0 19.6 10.1 15.7
*Water~ and solids~free basis.
··Viscosity measured on oil after coke was removed.
"-Run CBU-34 was run in the continuous unit. All other runs were performed in the batch autoclave.
For 10' API oil. 10 lbs salt/lOOO bbls is equivalent to 18 ppm Cl.
STRUCTURAL ANALYZES OF ZUATA HEAVY
CRUDE OIL FEEDS AND RUN PRODUCTS
(Wt %)
Structural data for the Zuata crude oil feed and product
is given in Table 4B.
TABLE4B
12.0 10.3 11.8
13.1 10.8 11.9
17.3 22.5 21.1
6.5 5.1 7.0
4.5 4.3 4.6
5.0 2.9 3.2
58.4 55.9 59.6
10-----TA-B-LE-4-B--con-tin-ue-d ------ STRUCTURAL ANALYZES OF ZUATA HEAVY
CRUDE OIL FEEDS AND RUN PRODUCTS
(Wt%)
6.3
4.5
2.3
0.6
4.3
29.2
11.2
ZU-4
100.0
ZU-1
5.9
4.9
2.6
1.3
_5_.6_
30.0
14.1
100.0
2.4
0.9
0.1
0.1
9.8
20.4
21.2
Feed
100.0
EXAMPLE S
Remainder
3-Ring Aromatics
4-Ring Aromatics
5-Ring Aromatics
Po1yaromatics
Sulfur Aromatics
Batch autoclave and continuous unit runs were conducted
on the Cerro Negro crude oil sample. The results
are given in Table SA.
TABLE SA
20
15
25
ZU-4
415
15
400
15
Feed ZU-1
Run Temperature, 'C.
Residence Time, Min.
Structure
Light Fractions
Paraffins
Cycioparaffins
Condensed
Cycioparaffins
Alkyl Benzenes
Benzo Cycioparaffins
Benzo Dicyloparaffins
Structure
Heavier Fractions
Run
Feed
CN 1
CN2
CN3
CN4
CN5
CBU-32
Run
Feed
CNI
CN2
CN 3
CN4
CN 5
CBU-32
CERRO NEGRO HEAVY OILS RUN DATA
Pres- Feed Product Viscosity" Residual Asphaltene' Solid
Temp sure, H2O Time H2O cp cp Gravity WI. Conv. WI. Alter. WI.
'C. psig % min··· % 25' C. 80' C. 'API % % % % %
9.8 321,000 1,780 8.0 65.5 21.8 0.37
350 1550 9.8 15 0.7 16,900 695 15.0 58.0 11.5 16.9 22.5 0.83
360 1525 9.8 15 2.3 11,500 402 12.7 54.7 16.5 18.1 16.9 0.10
370 1500 9.8 15 5.4 6,360 215 14.8 53.5 18.3 17.8 18.4 0.21
405 1630 9.8 15 2.6 5,150 159 14.3 53.8 17.9 18.4 22.9 1.01
415 1760 9.8 15 6.8 4,030 127 14.2 44.3 32.4 20.3 6.9 1.32
415 980 9.8 1.6 8.1 37,500 652 13.9 59.7 8.9 18.3 16.1 0.35
425 1030 9.8 1.4 5.8 13,600 352 12.5 56.0 14.5 18.2 16.5 0.42
435 1060 9.8 1.0 4.2 4,610 150 11.6 48.3 26.3 20.0 8.3 0.60
Coke Gas IBP- 450- Resid Con- Sulfur Volume %
Wt. WI. 450' F. 950' F. +950F Carbon WI. IBP-450' F. 450- 650- 450-950' F.
% % WI. % WI. % WI. % Wt.% %' Vol % 'API Sp gr 650' F. 950' F. 'API Sp gr
0.2 2.4 31.9 65.5 14.6 3.8 2.9 37.0 .840 11.7 22.8 19.5 .937
0.0 0.7 2.1 39.2 58.0 14.2 3.7 2.4 37.3 .838 18.1 22.4 19.8 .935
0.0 3.8 3.6 37.9 54.7 14.2 3.6 4.3 36.6 .842 18.7 21.1 19.8 .935
0.0 0.9 5.4 40.2 53.5 15.4 3.6 6.3 38.5 .832 21.4 20.0 19.4 .938
0.0 1.6 2.9 41.7 53.8 14.6 3.5 3.5 41.3 .819 18.4 25.2 20.8 .929
0.3 1.7 9.5 44.2 44.3 17.3 3.5 11.3 42.0 .816 23.0 22.8 19.2 .939
NO 2.4 4.0 . 33.8 59.7 15.3 3.3 4.7 35.1 .849 16.3 19.6 19.7 .936
ND 2.5 1.3 40.3 56.0 15.7 3.3 1.5 33.1 .860 19.5 23.7 20.2 .933
ND 7.9 2.6 41.1 48.3 15.6 3.3 3.2 36.8 .841 22.8 21.8 19.4 .938
Sulfur Distribution Pour
% % % Cl Point
Run Liquid Gas Solids ppm 'C. H2 Cf4
Feed 69.0 27
CNI 97 0 0 5.5 12
CN2 95 0 0 3
CN3 95 0 0 13.8 -1
CN4 93 0 0 9.2 4
CN 5 93 0 0
CBU-32 86 5 0 5 9.2 30.1
85 5 0 5 9.3 30.8
85 11 0 2 6.4 30.7
1.3
1.8
1.6
Gas Analysis, %
4.9 12.5 19.7 9.4
3.6 12.1 19.2 9.3
3.1 12.6 18.4 10.1
12.9
13.9
17.1
·Water4 and solids-free basis.
··Viscosity measured on oil after coke was removed.
"-Run CBU-32 was run in the continuous unit. All other runs were performed in the batch autoclave.
For 10' API oil. 10 lbs salt/lOOO bbls is equivalent to 18 ppm CI.
2-Ring Aromatics 7.1 9.7 11.2
Structural data for the Cerro Negro crude oil feed is
given in Table SB.
TABLE7A
Gravity Viscositv, cps
Wt,% Sp gr 'API 25' C. 80' C.
0.990 11.5 41,600 612
2.4 0.850 35.0 6 4
18.5 0.902 25.4 16 8
20.9 0.889 27.7 12 7
15.9 0.953 17.0 434 47
63.2 1.006 9.1 Solid Solid
79.1 0.998 10.2 Solid 17,700
VISCOSITY AND GRAVITY OF
COLD LAKE HEAVY OIL FRACTIONS
The whole oil and +650· F. fraction were then each
reacted in a series of bath rocking bomb autoclave experiments
at temperatures of 400· F. and 415· F. to
Whole Oil
-450
450-650
-650
25 650-950
+950
+650
Fraction
20 'F. ----....;........;.......:...:::....-_-----------
46
ing range between 450· F.-650· F.; (2) the +650· F.
primary fraction produced one fraction with a boiling
range between 650· F.-950· F., and one fraction with a
boiling range above 950· F. (+950· F.). In sum, the
5 produced fractions for testing were as follows:
-650· F. (primary fraction)
-450· F.
450· F.-650· F.
+ 650· F. (primary fraction)
650· F.-950· F.
+950· F.
The whole oil and the produced fractions were analyzed
and measured for weight (%), specific gravity,
•API, and viscosity (centipoise). The results are given in
15 Table 7A.
10
4,778,586
12.0
10.9
20.7
6.4
4.3
_7_.2_
61.5
12.1
2.1
0.9
0.2
0.1
-.2:.L
25.0
~
100.0
45
TABLE5B
EXAMPLE 6
Remainder
Structure
Heavier Fractions
2-Ring Aromatics
3-Ring Aromatics
4-Ring Aromatics
5-Ring Aromatics
Polyaromatics
Sulfur Aromatics
Structure
Light Fractions
Paraffins
Cycloparaffins
Condensed
Cycloparaffins
Alkyl Benzenes
Benzo Cycloparaffms
Benzo Dicycioparaffins
STRUCTURAL ANALYSES OF
CERRO NEGRO HEAVY CRUDE OIL FEED
(Wt%)
Batch autoclave runs were conducted on two shale
oil samples. The feed for Run OS-l was from the
Paraho Shale Oil operation. The feed for Runs as 4-6
were from another shale oil operation. The results are 30
given in Table 6A.
TABLE 6A
SHALE OIL ANALYTICAL RESULTS
Pres- Feed Product Viscosity Grav- Residual Asphaltene' Solid Coke Gas
Temp sure, H2O Time H2O cp cp ity Wt. Conv. Wt. Alter. Wt. WI. Wt.
Run 'C. psig % min % 25' C. 80' C. 'API % % % % % % %
Paraho Shale Oil - Batch Runs
Feed 0.0 0.0 Solid 24 21.8 22.9 1.8 0.02 0.07
OS-I 400 250 0.0 15 0.0 133 19 22,5 34.8 -52.0 3.2 -77.8 0.06 ND 2.0
Shale Oil - Batch Runs
Feed 2.4 552 9 23.1 12.7 2.0 0.34 1.0
OS-4 400 910 2.4 15 0.0 20 9 31.5 8.6 32.3 1.6 20.0 0.18 ND 2.1
OS-5 380 830 2.4 15 0.9 20 8 30.8 9.3 26.8 1.6 20.0 0.35 ND 2.0
OS-6 350 720 2.4 15 0.0 393 9 28.6 10.8 15.0 1.7 15.0 0.17 ND 0.7
IBP- 450- Resid Con- Sulfur' Volume % Pour
450' F. 950' F. +950' F. Carbon WI. IBP-450' F. 450- 650- 450-950' F. Point
Run WI. % Wt.% WI. % Wt.% % Vol % 'API Sp gr 650' F. 950' F. 'API Sp gr C.
Paraho Shale Oil - Batch Runs .
Feed 6.1 70.9 22.9 2.5 1.0
OS-I 5.3 57.8 34.9 4.6 0.8 6.1 22.5 .919 21.9 36.1 22.8 .917
Shale Oil - Batch Runs
Feed 5.5 80.7 12.7 2.6 2.1 6.1 40.8 .821 39.0 44.4 28.2 .886 20
OS-4 18.3 71.0 8.6 2.4 0.9 19.0 37.9 .835 41.4 27.5 26.8 .894 -1
OS-5 11.1 77.6 9.3 2.4 0.9 11.7 38.9 .830 44.7 31.4 27.7 .889 20
OS-6 12.3 76.2 10.8 1.8 0.9 13.1 38.6 .832 40.5 35.4 27.9 .888 20
·Water and solids free basis.
EXAMPLE 7
The Cold Lake heavy oil was distilled to produce
various fractions of different boiling point ranges. Ini- 60
tiaIly, the Cold Lake heavy oil was distilled to produce
two primary fractions: one fraction with a boiling range
of up to 650· F. (-650· F.) and one fraction with a
boiling range above 650· F. (+650· F.). Portions of
these two primary fractions were then further distilled 65
to give four additional fractions: (1) the -650· F. primary
fraction produced one fraction with a boiling
range of less than 450· F., and one fraction with a boilcompare
the effect of reaction temperature on viscosity
reduction in a whole oil fuel and a topped fuel. The
reaction times were 15 minutes. The temperature tests
produced a "whole oil product" and a "+650· F. product."
A portion of the +650· F. was blended with the
-650· F. fraction at the proportion of the original
whole oil to give a blended product. The viscosities of
the temperature reacted +650· F. fraction, the blended
product, and the temperature reacted whole oil were
measured and compared. Results are shown in Table
7B.
EXAMPLE 8
4,778,586
47 48
TABLE 7B
COMPARATIVE TEMPERATURE RUNS
As-
Resid phal-
Temp Time, Viscosity +950° F. tene Volume %
Run Feed 0e. min 25° C. We. Wt% Wt% 450° 450°_650° F. 650°-950' F.
I +650° F. 400 IS 7620 533 63.0 17.9 4.5 6.5 27.3
2 +650° F. 415 IS 1580 101 5\.5 19.4 10.9 13.9 25.0
+650° F. product from 400 IS 1330 57 49.8 14.1 6.0 23.6 2\.6
Run I, (400° C.), blended
with _650° F. fraction
+650° F. product from 415 IS 572 35 40.7 15.3 I \.0 29.4 19.8
Run 2, (415° C.), blended
with _650° F. fraction
3 Whole oil 405 IS 762 57 45.7 14.0 9.7 22.3 24.5
4 Whole oil 415 IS ISS 27 37.2 13.2 13.5 2\.9 26.9
mately 240 feet long with a 88-foot expanded section at
A run was made in a fifty barrel per day pilot plant, the bottom of the string. The expanded section was
designed to simulate operation in a larger scale vertical 20 2.62-inch LD. and gave approximately 15-minute retentube
reactor system. This run was performed to confirm tion time (based upon oil volume only) at a flow rate of
results obtained in the batch and continuous bench scale 1.5 gallon/minute. The reacted oil then flowed up the
experiments and to investigate heat transfer. The fol- i-inch center of the coaxial string. At the top of the
lowing is a description of the pilot plant: string the flow of product was through the k-inch cen-
An insulated and coiled truck tanker containing ap- 25 ter tube of the horizontal coaxial heat exchanger. Prodproximately
6,000 gallons of the heavy oil was located uct then flowed to the pressure letdown manifold which
adjacent to the test site. Steam was produced by a porta- directed the flow to either or both of the Greylok choke
ble boiler unit and circulated through the tanker coils to assemblies or bypassed the chokes and directed flow to
heat the oil to a temperature of approximately 1200 F. to a series of pressure letdown barstock valves.
1600 F. At this temperature, the oil was fluid enough to 30 The product then passed to the first gas-liquid separabe
circulated through the tanker by a Roper gear pump. tion tank. The liquid level in this tank was monitored by
Additionally, a 1,250-gallon heated and insulated tank a level indicator in order to maintain a liquid level in the
was provided for storage of feed oil and was also tank. The level was controlled by manually adjusting
equipped with a Roper gear pump and circulating loop. the liquid discharge valve on the bottom of the tank.
A bleed stream from either the trailer or circulating 35 This tank was kept at 10 to 25 psig to help the separation
loop supplied oil to either oftwo feed tanks. Exch of the of gas and liquid. The product was collected in a prodfeed
tanks was equipped with an Orberdorfer gear uct tank and transferred by pump into the product truck
pump and circulating loop. Each circulating loop had trailer except during product sampling periods.
two inline heaters, one on the pump inlet and one on the The gas flowed to the second phase separation tank
pump discharge, to heat the oil to 1650 F. to 1750 F. 40 where any light condensates were collected. Gas then
Each set of heaters had a temperature controller to flowed to the scrubber circuit through a gas meter, and
maintain the temperature of the oil in the tank. A bleed gas sampling loop.
stream from each of the feed tank circulating loops Gas flowed into the packed scrubber tower where it
supplied hot oil to the common suction manifold of the was contacted with a circulating 20% caustic (NaOH)
high pressure triplex pumps. All of the piping for the 45' solution spray. This solution removed the H2S from the
feed oil circuit was provided with temperature con- gas. The pH of this solution was monitored and fresh
trolled heat tape and fiberglass insulation. solution was pumped from the caustic makeup tanks
Two FMC Bean triplex piston pumps provided the into the scrubber tank to maintain pH. Both caustic
high working pressure of the system at flow rates of 1 makeup and waste solution removal were made with a
to 4 gpm. Only one of these pumps was in use at a time 50 variable speed dual head piston pump. The waste soluduring
actual operation; the second pump was a backup. tion was stored in appropriate tankage for treatment and
The high pressure discharge of each of these fed a com- disposal.
mon line to the coaxial heat exchanger. Also on the high A gas booster pump was used to pull the gas from the
pressure discharge ofthese pumps were Grear Pulsation scrubber circuit into the second section of the gas com-
Dampeners, pressure indicators, safety relief valves, 55 bustor unit where it was incinerated.
and rupture disks. The safety relief valves and rupture A Boscan, Venezuela crude was used as the feeddisks
had return lines to the feed tanks. stock. The pilot plant was operated for ninety-six hours,
High pressure feed oil was then pumped through the and 102.4 barrels of oil were processed at three condisurface
coaxial heat exchanger composed of a I-inch tions. Results are given in Table 8A. In the run 20 Ib of
diameter tube for the feed flow with a !-inch diameter 60 coke were produced, equivalent to 0.05 weight percent
tube inside carrying the product oil. The coaxial heat of the oil fed to the system.
exchanger flow can be configured to use two, four, or During this run, the reactor temperature (bulk fluid
all six sections of the heat exchanger unit. The heat temperature) was maintained at about 7500 F., 7600 F.,
exchanger was wrapped with temperature limiting 8 and 7650 F. as shown in Table 8B. The highest heater
watts/foot heat tape and fiberglass insulation. 65 temperatures measured were 777" F., 8040 F., and 8060
Feed flowed from the coaxial heat exchanger to the F. for these bulk fluid temperatures, giving the followouter
I-inch side of the I-inch by i-inch coaxial vertical ing ~T's: 270 F. (15 0 C.) @ 7500 F.; 440 F. (240 C.) @
geoclave reactor string. The I-inch string was approxi- 7600 F.; and 41 0 F. (23 0 C.) @ 7650 F.
49
4,778,586
50
TABLE8A
BOSCAN HEAVY OILS RUN DATA
Pres- Feed Product Viscosity"" Residual Asphallene" Solid
Temp sure, H2O Time H2O cp cp Gravity Wt. Conv. WI. Alter. WI.
Run 'C. psig % min··· % 25' C. 80' C. 'API % % % % %
Boscan Crude
Feed 1.2 57,957 828 9.5 64.1 19.0 0.12
Sample 1 395 1553 1.2 6.7 0.0 2,698 180 12.4 54.7 14.7 14.9 21.6 0.17
Sample 2 399 1594 1.2 6.1 0.0 2,095 131 12.6 56.6 11.7 15.5 18.4 0.21
Sample 3 399 2058 1.2 5.7 0.0 2,086 103 12.6 53.0 17.3 15.5 18.4 0.09
Sample 4 404 1995 1.2 7.1 0.0 1,085 64 12.9 50.4 21.4 15.8 16.8 0.08
Sample 5 408 2032 1.2 5.8 0.0 736 43 13.0 46.1 28.1 16.0 16.0 0.15
Sample 6 407 2088 1.2 4.8 0.1 857 50 13.2 47.5 25.9 15.8 16.8 0.11
Sample 7 407 2106 1.2 5.6 0.0 754 43 13.5 47.8 25.4 15.6 17.9 0.04
Sample 8 408 2071 1.2 5.8 0.0 934 46 13.2 46.7 27.2 15.8 16.8 0.11
Sample 9 406 2056 1.2 5.7 0.0 1,036 81 13.2 46.8 26.9 15.8 16.8 0.12
Sample 10 407 1982 1.2 5.3 0.1 842 55 13.5 48.4 24.6 15.6 17.9 0.14
Sample 11 404 2123 1.2 5.1 0.0 868 46 13.2 49.7 22.5 15.8 16.8 0.13
Sample 12 407 2000 1.2 4.5 0.0 1,137 58 13.0 48.1 24.9 15.7 17.4 0.17
Sample 13 408 2000 1.2 4.1 0.0 941 73 13.3 51.3 20.0 15.5 18.4 0.10
Sample 14 409 2124 1.2 3.1 0.1 1,123 67 13.2 51.3 19.9 15.7 17.4 0.12
Sample 15 406 2120 1.2 4.0 0.0 1,245 73 13.0 52.4 18.3 15.6 17.9 0.10
Sample 16 402 2007 1.2 4.1 0.0 989 66 12.9 50.7 20.9 15.7 17.4 0.11
Gas IBP- 450- Resid Con- Sulfur Pour IBP-450' F. Volume %
WI. 450' F. 950' F. +950F Carbon WI. Pt. Vol 450- 650- 450-950' F.
Run % WI. % WI. % WI. % Wt.%" % 'C. % 'API Sp gr 650' F. 950' F. 'API Sp gr
Boscan Crude
Feed 1.6 5.1 29.2 64.1 13.5 5.2 7 6.0 38.3 .833 18.0 13.2 21.6 .924
Sample 1 3.1 5.1 37.1 54.7 15.1 4.7 -5 6.0 36.6 .842 19.0 20.9 21.3 .926
Sample 2 2.6 5.7 35.1 56.6 14.9 4.8 -12 6.8 40.0 .825 16.2 21.2 21.8 .923
Sample 3 4.7 6.2 36.2 53.0 14.6 4.8 -12 7.3 36.9 .840 16.5 22.1 21.1 .927
Sample 4 3.0 8.3 38.4 50.4 15.5 4.4 -15 9.6 36.2 .844 20.0 20.6 21.8 .929
Sample 5 3.0 8.6 42.4 46.1 15.9 4.5 -19 10.2 37.2 .839 19.4 19.9 21.3 .926
Sample 6 4.9 9.2 38.5 47.5 15.3 4.5 -22 10.9 38.3 .833 19.2 21.9 20.8 .929
Sample 7 6.6 5.4 40.2 47.8 15.9 4.4 -21 6.4 38.1 .835 17.4 25.9 21.1 .927
Sample 8 4.7 11.2 37.5 46.7 15.1 4.4 -16 13.1 36.1 .844 17.0 22.7 19.8 .935
Sample 9 4.0 9.2 40.0 46.8 16.0 4.5 -17 11.0 38.7 .831 21.3 21.3 20.5 .931
Sample 10 4.6 7.1 40.0 48.4 15.1 4.4 -17. 6.8 41.1 .820 20.7 20.1 21.8 .923
Sample 11 4.2 6.6 39.6 49.7 13.6 4.6 -18 7.9 38.6 .832 21.1 21.5 21.5 .925
Sample 12 3.7 11.3 36.9 48.1 15.4 4.5 -18 13.5 37.4 .838 18.5 21.2 20.5 .931
Sample 13 3.9 7.1 37.7 51.3 14.8 4.6 -18 8.5 39.7 .827 19.8 20.8 21.6 .924
Sample 14 4.0 7.6 37.1 51.3 16.0 4.6 -18 9.2 40.4 .823 19.6 20.4 21.5 .925
Sample 15 2.6 - 6.7 38.3 52.4 15.5 4.5 -15 8.0 39.7 .826 19.7 21.4 21.8 .923
Sample 16 2.4 7.6 39.2 50.7 15.8 4.4 -14 9.1 39.7 .827 19.3 22.6 21.5 .925
Sulfur Distribution
% % % Gas Analysis, %
Run Liquid Gas Solids H2 CH4 CO CO2 C2H6 H2S C3Hg C2H4 C3H6 Other
Boscan Crude
Feed
Sample 1 89 9 0 3.6 26.4 0.5 4.2 11.2 32.2 7.7 0.2 1.8 10.9
Sample 2 92 4 0 1.8 25.4 0.3 4.6 11.4 33.2 8.1 0.2 1.8 11.8
Sample 3 90 10 0 1.8 25.9 0.3 4.1 11.7 33.3 8.1 0.2 1.7 11.0
Sample 4 84 5 0 1.8 29.8 0.1 4.0 11.9 31.3 8.1 0.1 1.3 11.5
Sample 5 85 10 0 1.7 26.8 0.2 3.2 11.3 36.7 8.1 0.1 1.1 10.8
Sample 6 85 13 0 1.8 28.5 0.0 3.8 12.3 31.0 8.5 0.1 1.2 12.9
Sample 7 82 15 0 1.8 28.2 0.1 3.7 12.5 31.6 9.2 0.1 1.0 11.8
Sample 8 83 14 0 1.4 30.0 0.0 3.8 12.8 30.9 9.0 0.1 1.1 10.8
Sample 9 85 13 0 0.8 30.2 0.2 3.1 13.2 31.0 9.3 0.1 1.3 10.8
Sample 10 84 15 0 1.6 25.6 0.0 3.2 11.0 38.9 8.0 0.1 1.1 10.6
Sample 11 86 12 0 1.9 31.9 0.2 3.7 12.9 30.3 8.6 0.1 1.1 9.4
Sample 12 85 14 0 1.3 31.0 0.1 3.2 11.2 29.7 14.2 0.1 0.9 8.2
Sample 13 86 15 0 1.1 30.0 0.6 3.5 12.7 31.1 8.7 0.1 0.7 10.9
Sample 14 86 16 0 0.7 29.9 0.1 3.4 13.0 32.5 9.0 0.1 1.1 10.3
Sample 15 86 6 0 0.8 30.4 0.2 3.5 12.9 32.4 9.0 0.1 1.2 9.6
Sample 16 83 8 0 1.5 29.6 0.0 3.4 12.8 30.6 9.2 0.1 1.3 11.6
·Water~ and solids~free basis.
"'Viscosity measured on oil after coke was removed.
·"'Residence time for continuous unit was calculated for temperatures within 50 C. of reaction temperature.
TABLE8B TABLE 8B-continued
Sample (1) Reactor Temp., 'F. (2) Heater Temp., 'F. Sample (1) Reactor Temp., 'F. (2) Heater Temp., OF.
# Top Bottom (3) Top (3) Bottom # Top Bottom (3) Top (3) Bottom
65
1 745 743 764 752 5 766 767 794 788
2 747 750 777 763 6 763 764 804 797
3 748 750 778 765 7 764 764 802 797
4 758 759 788 779 8 767 766 799 791
4,778,586
798 790
802 797
791 787
804 801
806 804
796 792
779 772
770 763
(2) Heater Temp.• 'F.
(3) Top (3) Bottom
51
EXAMPLE 9
TABLE 8B-continued
763 763
764 765
759 760
764 765
764 766
765 768
761 762
760 756
(I) Reactor Temp.• 'F.
Top Bottom
9
10
11
12
13
14
15
16
Sample
#
52
8.5 million BTU!hr is used to heat the heat exchange
fluid.
The crude oil feed stream which has been heated to
about 375° C. and whose pressure has increased from an
5 inlet pressure of 50 psig to a pressure of about 1500 psig
enters the outer reactor pipe. The temperature of the
stream is increased to a reaction temperature of about
400° C. The pressure is increased to about 1750 psig.
The temperature differential between the bulk tempera-
10 ture of the hydrocarbon stream and the heat exchange
fluid is less than 25° C. The hydrocarbon stream passes
(I) Bulk temperature of fluid measured at top aod bottom of the lower 22 feet of through the outer reactor pipe and into the inner reac-
~~::~~~:·With thermocouple adjacent to heater. tor pipe at a flow rate which provides a total reactor
(3) Heater located within one foot of top aod bottom of lower 22 feet of reactor residence time of about 12 minutes at a hydrocarbon
string. 15 stream feed rate of 10,000 barrels per day. As the processed
hydrocarbon stream passes out of the inner reactor
pipe and into the riser pipe, cooling of the processes
stream is initiated by heat exchange contact with the
incoming hydrocarbon feed stream. The temperature
and pressure of the processed stream decreases as it
flows upward from the reactor zone. When the processed
stream exits the riser pipe the temperature is
about 125° C. and the pressure is about 250 psig.
Upon leaving the reactor system the process stream is
fed into a depropanizer in which the primary product is
separated from propane, water, and other gases. This
gas stream which amounts to about 1 million standard
cubic feet per day is further processed in a sequential
process stream to recover sulfur, process fuel, and natural
gas in an environmentally acceptable manner. The
primary product, which now has a viscosity of about
1000 cps at 25° c., is then introduced back into a transportation
network for transport to a refinery or transshipment
point.
While various embodiments of the present invention
have been described in detail, it is apparent that modifications
and adaptations of those embodiments will
occur to those skilled in the art. However, it is to be
expressly understood that such modifications and adaptations
are within the spirit and scope of the present
invention, as set forth in the following claims.
What is claimed is:
1. A method for improving the transportability of
hydrocarbons said method comprising:
(a) flowing an influent hydrocarbon feed stream at a
first temperature and a first pressure into a downcomer
to form a hydrostatic pressure head and
provide a pressurized feed stream at a second pressure;
(b) heating said influent stream by heat exchange with
an effluent treated hydrocarbon stream wherein at
least one of said streams is in turbulent flow to
increase the temperature of said influent stream
from said first temperature to a second temperature
and provide a heated feed stream;
(c) contacting said heated and pressurized feed stream
with an active heat source in a reaction zone to
provide the feed stream at a reaction temperature
between about 300° C. and the coking temperature
of said hydrocarbons and a reaction pressure of at
least about 1000 psi to form said treated hydrocarbon
stream;
(d) maintaining a temperature differential between
said active heat source and said feed stream in said
reaction zone of less than about 30° C. to form a
treated hydrocarbon stream; and
(e) removing said treated stream from said reaction
zone by passing said treated stream upward in a
A heavy crude oil having a viscosity in excess of
200,000 cps is passed through a dewatering process to 20
reduce the basic sediment and water (BSW) of the produced
oil to less than 5 weight percent. The resulting oil
is then passed into storage tanks. For convenience the
storage tanks are sized to provide at least a 24 hr supply
of feed oil at a use rate of 10,000 barrels per day. The 25
treated oil is then passed from the storage system or
alternatively directly from the BSW unit to the processing
unit. This processing unit is located in a vertical
shaft having a depth of about 4,500 ft and a finished
casing diameter of 24 in. Suspended in the vertical shaft 30
is the reactor string which consists of two concentrically
oriented pipes which comprise adowncomer-riser
system. Attached to the bottom of the downcomer-riser
system is the reactor which consists of an inner reactor
pipe and an outer reactor pipe. The downcomer pipe is 35
a 14 in. diameter pipe. The riser pipe which is located
inside the downcomer is 10 in. diameter. The outer
reactor pipe has a 20 in. diameter and is 464 ft in length.
The inner reactor pipe, which is located within the
outer reactor pipe, is 464 ft in length with a 10 in. diame- 40
ter. The inner and outer reactor pipes together comprise
a reactor volume of 880 cubic ft which provides a 12 to
15 min residence time at reaction temperature and pressure
with about a 2 weight percent steam and about 2
weight percent gas content of the hydrocarbon stream. 45
The crude oil feed enters the reactor string at about
60° C. to about 100° C. and travels downward through
the annular portion of the concentric pipe downcomerriser
system. The oil is heated through indirect heat
exchange with processed oil which is traveling upward 50
in the center riser pipe. The crude oil stream is heated to
within 25° C. of the reaction temperature before it enters
the outer reactor pipe. Supplemental heat is supplied
by means of indirect heat exchange with a hightemperature
pressure-balance fluid which occupies the 55
void volume surrounding the reactor string. With a 25°
C. approach temperature at the hot end of the riser
downcomer heat exchanger, the system heat duty is
about 5.64 million BTU/hr. In order to account for
well-casing heat losses, this value is increased by 50 60
percent to 8.46 million BTU/hr. A heat exchange fluid
flow rate of 1,060 gal/min is required to supply this heat
duty at a hot fluid-reactor approach temperature of 25°
C. The heat transfer fluid is circulated via a 3 in. pipe
using a 50 psi high-temperature centrifugal pump. A gas 65
cap is maintained above the heat exchange fluid to provide
the primary pressure drive forced to overcome the
pressure head. A surface gas-fired tube heater rated at
4,778,586
53
riser to form said effluent treated stream of reduced
viscosity.
2. The method of claim 1 wherein said reaction pressure
is between about 1000 and about 4000 psi.
3. The method of claim 2 wherein said reaction tem- 5
perature is between about 350° C. and about 475° C.
4. The method of claim 2 wherein said reaction temperature
is between about 375° C. and about 435° C.
5. The method of claim 1 wherein said contacting
with said active heat source provides a coke make of 10
less than about 0.5 weight percent of said hydrocarbon
stream.
6. The method of claim 1 wherein said turbulent flow
is multiphase flow.
7. The method of claim 6 wherein said influent stream 15
and said effluent stream are each in multiphase flow.
8. The method of claim 1 wherein said temperature
differential is less than about 15° C.
9. The method of claim 1 wherein said temperature
differential is less than about 5° C. 20
10. The method of claim 1 wherein said hydrocarbon
feed stream is selected from the group consisting of
whole crude oil, kerogen, bitumen, shale oil, tar sands
oil, and mixtures thereof.
11. The method of claim 1 wherein said hydrocarbon 25
feed stream has an initial API gravity at 25° C. below
about 20° and said treated hydrocarbon stream has an
API gravity at least 2° higher than that of said hydrocarbon
feed stream.
12. The method of claim 1 wherein said first pressure 30
is less than about 500 psi.
13. The method of claim 1 wherein said treated hydrocarbon
stream is removed from said riser and gaseous
materials are separated from said stream.
14. The method of claim 1 wherein said treated hy- 35
drocarbon stream is removed from said riser and a portion
of components boiling below about 40° C. are separated
from said treated stream and introduced into said
hydrocarbon feed stream.
15. The method of claim 1 wherein said first tempera- 40
ture is less than about 100° C. said first pressure is less
than about 200 psi said reaction temperature is between
about 350° C. and about 450° C. said reaction pressure is
between about 1000 psi and about 2000 psi said second
temperature is above about 250° C. and said tempera- 45
ture differential is less than about 25° C.
16. The method of claim 1 wherein said hydrocarbon
feed stream comprises up to about 10 weight percent
water.
17. The method of claim 1 wherein said treated hy- 50
drocarbon stream is removed from said riser and
blended with untreated hydrocarbon.
18. The method of claim 1 wherein said hydrocarbon
feed stream consists essentially of a heavy oil, water and
a diluent wherein said water is present in an amount less 55
than about 10 weight percent of said feed, and said
diluent is a light fraction of hydrocarbons which is
present in an amount sufficient to render said heavy oil
pumpable.
19. The method of claim 18 wherein said heavy oil is 60
whole crude oil.
65
54
20. A method for decreasing the viscosity of hydrocarbons
said method comprising:
(a) providing an influent hydrocarbon feed stream at
a temperature TI and a pressure PI;
(b) passing said influent stream downward in a downcomer
to form a hydrostatic pressure head and
increase pressure on said influent stream to provide
a pressurized feed stream;
(c) heating said influent stream by heat exchange
contact with an effluent stream wherein said
streams are in multiphase flow to increase the temperature
of said influent stream from temperature
T, to temperature T2, which is within about 50° C.
of a reaction temperature and provide a heated feed
stream;
(d) contacting said heated and pressurized feed
stream with an active heat source having a temperature
differential between said heat source of said
feed stream of less than about 30° C. in a reaction
zone to provide the feed stream at a reaction temperature
of between about 300° C. and the coking
temperature of said hydrocarbons and a reaction
pressure of at least about 1000 psi;
(e) maintaining said feed stream in said reaction zone
to reduce the viscosity of said feed stream and form
a treated hydrocarbon stream; and
(t) removing said treated stream from said reaction
zone and passing it upward as said effluent stream
in a riser into said heat exchange contact with said
influent stream.
21. The method of claim 20 wherein said effluent
stream is removed from said riser and at least a portion
of components boiling below about 40° C. are separated
from said stream and are introduced into said influent
feed stream.
22. The method of claim 20 wherein said effluent
stream is removed from said riser and is blended with
untreated heavy oil to reduce the viscosity of said heavy
oil.
23. The method of claim 20 wherein said hydrocarbon
feed is selected from the group consisting of whole
crude oil, bitumen, kerogen, shale oil, tar sands oil, and
mixtures thereof.
24. The method of claim 20 wherein said reaction
temperature is between about 350° C. and about 475° C.
and said reaction pressure is between about 1000 psi and
about 2000 psi.
25. The method of claim 20 wherein said reaction
pressure is between about 1000 psi and 4000 psi.
26. The method of claim 20 wherein said hydrocarbon
feed consist essentially of a heavy oil, water and a
diluent, wherein said water is present in an amount less
than about 10 weight percent of said feed and said diluent
is a light fraction of hydrocarbons which is present
in an amount sufficient to render said heavy oil pumpable.
27. The method of claim 26 wherein said heavy oil is
whole crude oil.
28. The method of claim 20 wherein said temperature
differential is less than about 15° C.
* * * * *
UNITED STATES PATENT AND TRADEMARK OFFICE
CERTIFICATE OF CORRECTION
PATENT NO.
DATED
INVENTOR(S)
4,778,586
October 18, 1988
Bain et al.
It is certified that error appears in the above-identified patent and that said Letters Patent
is hereby corrected as shown below:
Column la, line 7, please delete the letter liS" from
the word hours.
Column 19, Table lC, in the first line of the headings
please delete IIlPB II and insert -- lBP -- therefor.
Column 47, line 36, please delete "Exch" and insert
-- Each -- therefor.
Note:
Column 32, line 33, is a continuation of Table 2B.
Column 43, last line of the column, should follow the
term "Heavier Fractions ll in Table 4B.
Signed and Sealed this
Seventh Day of l\'Iarch, 1989
Attest:
DONALD J. QUIGG
Attesting Officer Commi.\'sioner of Patents and Trademarks