4,778,586 Viscosity reduction processing at elevated pressure

Patent Number/Link:

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

j~34

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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

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

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-

(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

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.

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

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

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·

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

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

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

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

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.

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

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

TABLEIB

MASS SPECTROMETER ANALYSIS OF

285-430' F. FRACTION OF THE COLD LAKE FEED

Paraffins 35.3 vol %

Olefins NO

4,778,586

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.

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

25.2

(3)

0.39

149

9.34

7.4

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.

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

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

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

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

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

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

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

4,778,586

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

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 %

35.3

12.7

100.0 vol %

4,778,586

TABLE 2B-continued

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

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

42.4

12.0

27.44

42.57

26.11

38.96

18.6

0.9424

4.54

0.16

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.

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

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

*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.

Run

Feed

BO I

BO 2

CBU-7

Run

. Feed

BO I

B02

CBU-7

4,778,586

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

-!!.:.!...

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

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

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

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

4,778,586

TABLE 2G-continued

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

_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

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

4,778,586

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

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

ZU-4

415

400

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. ----....;........;.......:...:::....-_-----------

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.

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

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.

4,778,586

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

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

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

Sample

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;

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

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.

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;

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