Process scheme for sequentially hydrotreating-hydrocracking diesel and vacuum gas oil

BACKGROUND OF THE INVENTION

The invention relates to hydrotreating and hydrocracking vacuum gas oil, Diesel, kerosene and naphtha fractions obtained from a vacuum residue hydrocracking (VRHCK,) using a high temperature and high pressure stripping and washing stage for connecting two high pressure reaction zones. More particularly, the invention relates to a hydrotreating and/or hydrocracking process using a high pressure, high temperature stripping and washing process which is well suited as an intermediate step integrated into a high temperature and high pressure loop of a vacuum residue hydrocracking (VRHCK) process wherein the VRHCK process and stripping-washing step produce kerosene, naphtha, Diesel and vacuum gas oil feeds to a hydrotreating-hydrocracking reactor for reducing sulfur content and improving naphtha, kerosene and Diesel yields and quality; in particular to produce a high smoke point kerosene, high cetane number and low sulfur Diesel and a low sulfur naphtha fractions.

Many refineries hydrotreat virgin and cracked feedstocks in order to obtain upgraded gasoline and Diesel products. These refineries utilize high-pressure units. High pressure hydrodesulfurization (HDS) units can be utilized with cracked vacuum gas oil (VGO), and when operated between 700-1200 psig, can achieve HDS conversion rates of greater than 99% so as provide a product having a sulfur content between 0.002 and 0.12% wt. This product can then be fed to a fluid catalytic cracking (FCC) process to produce gasoline and Diesel fuel with sulfur content less than 150 ppm and 600 ppm respectively. Unfortunately, the Diesel fraction produced in an FCC process from such a VGO feed typically has a cetane number of only about 20-30, which is considered “out of spec” and prevents this product from being incorporated into the Diesel pools. In order to be used, this Diesel fraction must be treated with additional hydrotreating steps. In addition, numerous other Diesel streams are readily available in the refineries such as straight run kerosene and Diesel, thermal cracked Diesel and the like, all of which have high sulfur content and typically medium cetane number that will require an additional deep hydrotreatment.

Conventional low-medium pressure Diesel hydrotreatment can satisfactorily reduce the sulfur content, but provides only small improvements in cetane number, in the range of 2-4 point increments.

Some refineries have installed residue hydrocracking units to convert the heaviest part of the oil into distillates. All of these technologies produce a medium quality product that requires an additional high pressure hydrotreating stage in a separate unit to fulfill the pool specifications. They also produce a large amount of VGO that needs to be converted in standard fluid catalytic cracking or hydrocracking units.

Typical catalysts for use in hydrotreating-hydrocracking processes to increase cetane number and smoke point and to reduce sulfur to very low level are extremely sensitive to even small amounts of sulfur and ammonia, and therefore cannot readily be incorporated into hydrotreating-hydrocracking reactors.

Alternatives for processing all the naphtha, kerosene, Diesel and VGO to solve the yield and quality problem identified above include integrating a high pressure hydrotreating and hydrocracking reactor into an existing or new vacuum residue hydrocracking technology for sequentially hydroprocessing-hydrocracking the available “out of spec” products. That effort attempts to address the yield, sulfur, smoke point and cetane number objectives by using two or more stages of hydroprocessing. Unfortunately, conventional vacuum residue hydrocracking units require a conventional separation process where the products are cooled and depressurized to be able to separate and then treat those products in a stand alone hydrotreating or hydrocracking or FCC unit. These conventional separation processes are carried out at low temperature, low pressure, or both, resulting in the need for additional compression systems, one for each stage, which can double equipment and operation costs. Even then, a complete and potentially very expensive hydrotreating-hydrocracking-FCC unit is required.

It is clear that the need remains for a method for treating feedstocks such as vacuum gas oil, kerosene, Diesel and naphtha, “out of spec” product from primary conversion (VRHCK), and others such as Diesel and naphtha fractions available in the refinery, so as to advantageously improve yields and reduce sulfur content while improving smoke point and cetane number. Further, the need remains for a process whereby separation of components is achieved at high temperature and pressure so as to avoid the need for additional compression equipment and the like. Still further, the need remains for a process to hydrofinish the product without the need to buy additional complete hydrotreating-hydrocracking units.

It is therefore the primary object of the present invention to provide a process whereby VGO, kerosene, Diesel, and naphtha “out of spec” feedstocks can advantageously and economically be converted into valuable end products.

It is another object of the invention to provide a process which can advantageously find use in revamping existing facilities or building new ones.

It is a further object of the invention to provide a process for high pressure and high temperature separation to produce an intermediate feedstock that can be blended with an external VGO, Diesel and naphtha “out of spec” component to be sequentially hydrotreated or hydrocracked in subsequent hydrotreating-hydrocracking reaction stages.

It is still another object of the present invention to utilize the existing high pressure and temperature loop in a vacuum residue hydrocracking unit or process to provide an integrated process which sequentially integrates a hot separation-stripping washing system into a vacuum residue hydrocracking process with a further hydrotreating-hydrocracking stage or stages so as to advantageously and efficiently provide for excellent conversion rates in converting the VGO, kerosene, Diesel and naphtha feeds to desirable end product, while nevertheless minimizing expense due to equipment costs and efficiently utilizing high pressure and temperature conditions.

Other objects and advantages will appear herein below.

SUMMARY OF THE INVENTION

In accordance with the present invention, the foregoing objects and advantages have been readily attained.

According to the invention, a process is provided for sequentially hydrotreating vacuum gas oil, kerosene, Diesel and naphtha, which process comprises the steps of providing a reaction feed containing residue, vacuum gas oil, kerosene, naphtha, Diesel, hydrogen sulfide, ammonia, and C

1

-C

4

gas phase compounds; providing a stripping gas; providing a washing feed; and feeding said reaction feed, said stripping gas and said washing feed to a stripping and washing zone so as to obtain a gas phase containing said hydrogen sulfide, said ammonia, said C

1

-C

4

gas phase compounds, said naphtha, said kerosene, said Diesel and said vacuum gas oil and a liquid phase, wherein said reaction feed is provided at a reaction feed pressure of between about 700 psig and about 3500 psig, and wherein said stripping and washing zone is operated at a pressure within about 80 psig of said reaction feed pressure. The hydrotreating-hydrocracking reactors, as well as the stripping/washing separator, are advantageously operated at substantially the same pressure, and preferably substantially the same temperature, thereby avoiding the need for additional compressor equipment between stages and limiting the need for additional heating between stages as well, while utilizing the already established temperature and pressure from the VRHCK loop to carry out further improvement as desired.

The inventive separation-reaction process can advantageously be used to integrate hydrotreatment-hydrocracking processes which occur at high temperatures and pressures so as to sequentially treat vacuum gas oil, kerosene, Diesel, naphtha and other feeds to obtain high conversion to desirable products at reduced cost.

BRIEF DESCRIPTION OF THE DRAWINGS

A detailed description of preferred embodiments of the present invention follows, with reference to the attached drawings, wherein:

FIG. 1

schematically illustrates a system and process in accordance with the present invention;

FIG. 2

further illustrates the integration of

FIG. 1

in a vacuum residue hydrocracking unit;

FIG. 3

illustrates stripping and washing steps in accordance with one embodiment of the invention;

FIG. 4

illustrates stripping and washing steps in accordance with another embodiment of the invention;

FIG. 5

illustrates still another embodiment of stripping and washing steps of the present invention;

FIG. 6

schematically illustrates the integration of a two stage hot separation washing and stripping in a vacuum residue hydrocracking process in accordance with the present invention;

FIG. 7

illustrates an integrated process in accordance with the present invention which is oriented to naphtha production;

FIG. 8

illustrates an integrated process in accordance with the present invention which is oriented toward production of high quality Diesel;

FIG. 9

illustrates an integrated process in accordance with the present invention wherein the system is oriented to hydrocracking for production of jet fuel and Diesel fuel;

FIG. 10

illustrates a basic configuration of an integrated process in accordance with the present invention for hydrodesulfurization of VGO and mild hydrocracking;

FIG. 11

presents a mass and energy balance example of HSSW in accordance with the invention; and

FIGS. 12 and 13

comparatively illustrate a separate conventional process (

FIG. 12

) and an integrated process in accordance with the present invention (FIG.

13

).

DETAILED DESCRIPTION

The invention relates to a process for sequentially treating vacuum gas oil, kerosene, Diesel and naphtha so as to provide a final product fraction including components having satisfactorily low sulfur content, kerosene and/or jet fuel fraction with high smoke point and Diesel fractions having cetane numbers sufficiently improved to allow incorporation into the Diesel pools. The process utilizes a stripping and washing step (HSSW) to accomplish a high temperature and high pressure separation of an intermediate feedstock for example from a vacuum residue hydrocracking process, so as to avoid the need for intermediate compression and/or reheating of the separated phases and thereby allow direct feed to subsequent hydrotreating-hydrocracking stages and the like which utilize the high pressure and temperature.

The invention further relates to an integrated processing method which integrates a hydrotreating-hydrocracking stage or system into a vacuum residue hydrocracking process having a high-pressure and high-temperature circuit and utilizing the aforesaid stripping and washing step between stages.

As will be further discussed below, the process of the present invention advantageously maintains the pressure of the product of the initial hydrotreating or hydrocracking step through separation of that product into portions, and through feed of some portions into a subsequent step such as a hydrotreating-hydrocracking step so as to provide the desired hydrotreating-hydrocracking conditions and reactions without the need for multiple compressors and the like, and to provide more efficient energy utilization. Conventionally, the intermediate feed, for example from a VGO reactor product, is cooled and the pressure reduced, to provide a separate hydrogen rich phase and a hydrocarbon rich phase. This creates the need for additional compressors and/or heating equipment to re-pressurize and re-heat at least some portions of the intermediate feed.

One process in which a hot separation-stripping and washing (HSSW) step according to the present invention is particularly advantageous is a process for sequentially treating a combined vacuum gas oil (VGO), kerosene, Diesel and naphtha feedstock. In such a process, the initial feed composed of hydrogen and vacuum residue is hydrocracked into a reaction feed containing VGO, Diesel, naphtha, C

1

-C

4

, hydrogen sulfide, ammonia and hydrogen. This reaction feed is preferably treated under high pressure and high temperature conditions utilizing a hot separation-stripping and washing (HSSW) zone as discussed further below, so as to obtain a gas phase which can advantageously be passed to a hydrotreatment-hydrocracking zone and a liquid phase which may suitably be treated through distillation to recover vacuum gas oil to be recycled to the process and a bottom that is ideal for feed to delayed coking processes or the like. The following description will be given in terms of this type of process. It should readily be appreciated, however, that the intermediate hot separation, stripping and washing steps and process of the present invention would be readily applicable to other types of processes as well and can be varied without departing from the scope of the present invention.

Typical feed for the overall process of this embodiment includes vacuum residue and various distillate products, one suitable example of which is vacuum gas oil (VGO), Diesel (DO) and naphtha (N). VGO streams are readily available in refineries but frequently have unacceptably high sulfur content. These streams do include portions which can advantageously be converted into useful gasoline, jet fuel and Diesel fractions in a stand alone hydrocracker, but at great expense. The VGO can also be treated in a Fluid Catalytic Cracking process but, unfortunately, the Diesel fraction produced typically has a cetane number which is too low to be useful without further treatment.

Additional feedstocks which can find advantageous use in the overall process of the present invention include other refinery streams including kerosene, Diesel and naphtha streams such as straight run kerosene and Diesel, thermal cracked Diesel and naphtha (for example from a delayed coker) and the like, each of which typically has high sulfur content, and Diesel having a medium cetane number which will require improvement in order to be added to the Diesel pool.

In accordance with the process of the present invention, a first hydrocracking reaction zone is established, for advantageously converting the vacuum residue into a product stream containing residue, VGO, kerosene, Diesel, naphtha, C

1

-C

4

gas phase compounds, hydrogen sulfide, and ammonia. Product fractions from the hydrocracking are used as reaction feed to a high pressure separation-stripping-washing (HSSW) zone operating at substantially the same pressure as the outlet from the hydrocracking step. The hot separation-stripping and washing step advantageously produces a gas phase advantageously containing hydrogen, naphtha, kerosene, Diesel, light and heavy vacuum gas oil, C

1

-C

4

containing compounds, H

2

S and NH

3

fractions, and a liquid phase including small amounts of Diesel and light vacuum gas oil, heavy vacuum gas oil and residue. The gas phase is advantageously still at a pressure and temperature which is sufficiently high that the gas phase can be fed directly to a second high pressure reaction zone, for example hydrotreating-hydrocracking, to convert the vacuum gas oil into high quality distillates and improve the quality of the Diesel, kerosene and naphtha, without the need for additional compressors or heaters and the like.

Thus, the stripping and washing step of the present invention provides the desired liquid and gas phases without substantially changing the pressure between the hydrocracking and subsequent hydrotreating-hydrocracking steps. The pressure at the hydrocracking or first stage, the separating stage and the hydrotreating-hydrocracking or second stage may advantageously be between about 900 psig and about 3500 psig, more preferably between bout 1300 psig and about 3000 psig. The pressure is preferably between about 900 psig and about 3000 psig at the outlet of the vacuum residue hydrocracking stage, and is maintained within about 80 psig, more preferably about 50 psig, of the pressure of the second hydrotreating-hydrocracking stage reaction inlet through the hot separation stripping and washing zone.

As set forth above, the feed to the hydrocracking reactor is preferably a vacuum residue oil feed that is to be converted to more valuable distillates products. The vacuum residue feed may be heated before entering the hydrocracking reactor, preferably to a temperature of between about 400° F. and about 800° F., and more preferably between about 500° F. and about 650° F. The vacuum residue may be blended with hydrogen and catalyst and fed to the hydrocracking reactor. The hydrocracking reactor may suitably be an up flow bubble type reactor or other suitable reactor, and the catalyst used may be any iron based or coke based catalyst having activity toward the desired reaction.

The product of the hydrocracking reactor typically includes hydrogen, naphtha, kerosene, Diesel, LVGO, HVGO, C

1

-C

4

hydrocarbons, hydrogen sulfide and ammonia and where the Diesel, kerosene and naphtha fractions are frequently “out of spec”, or have inappropriate characteristics for final blending in the refinery pool.

As used herein, the term “out of spec” refers to a hydrocarbon product or composition having characteristics which do not meet standards set for use of such a composition as a fuel.

This product stream, or at least a portion of the stream, is fed alone or combined with external VGO, Diesel, and/or naphtha, as a reaction feed to the high temperature and high pressure hot separation-stripping-washing zone for separation into phases as desired in accordance with the invention. An embodiment of a high pressure hot separation-stripping washing zone is illustrated in

FIG. 1

, which will be further discussed below.

FIG. 1

shows a hot separation-stripping-washing zone

10

, wherein the reaction feed

12

from the hydrocracking reactor is preferably introduced into a stripping and washing (HSSW) vessel or vessels

14

,

15

along with a stripping gas

16

such as hydrogen and a washing feed or medium

18

such as additional external feed of Diesel, LVGO and the like. A portion of washing feed in

FIG. 1

is shown passing from an external VGO source

19

through a furnace

20

and then being blended with a gas phase from vessel

15

, while another portion is fed to vessel

15

. Ideally, the reaction feed, washing feed and stripping gas are fed to the reactor or vessel each at different vertical heights, and the HSSW vessels

14

,

15

have a gas phase outlet

21

and a liquid phase outlet

22

.

The stripping gas serves to enhance high temperature and high pressure separation of hydrogen, hydrogen sulfide, ammonia, C

1-

C

4

and naphtha compounds into the gas phase. The hydrogen stripping also serves to enhance separation of the gas phase, and is itself present in the gas phase which is produced and which is useful as a feed to later hydrotreatment processes. In the vacuum residue hydrocracking (VRHCK) example of the present embodiment, the gas phase product of the stripping and washing step preferably includes hydrogen, naphtha, kerosene, Diesel, LVGO, C

1-C

4

hydrocarbons, H

2

S and NH

3

. The remaining liquid phase from the stripping and washing step mainly includes solid catalyst from the VRHCK process, HVGO and unconverted residue.

The liquid phase can advantageously be fed to further fractionation to recover VGO, which can be recycled, and catalyst plus unconverted residue that can be further processed, for example in a delayed coking process or the like. In the VRHCK embodiment of the present invention, this liquid phase may typically include a small amount of Diesel and LVGO, as well as HVGO and vacuum residue. A significant benefit provided by the present invention is that the washing step controls and/or avoids asphaltene and catalyst carry over into the gas phase which is beneficial for allowing further trickle bed processing.

It should readily be appreciated that the hot separation-stripping and washing steps of the present invention provide for advantageous separating of the gas and liquid phases, and the components present in each, without cooling and de-pressurization of the reaction feed which therefore does not require re-pressurization in order to be treated in subsequent high-pressure reactions.

It should also be noted that the use of externally obtained feed as washing and/or stripping feed allows for the adjustment or fine-tuning of temperature in the stripping and washing (HSSW) vessel, if desired. This is accomplished by feeding the external feed and/or stripping gas in greater or lesser amounts, and/or at different temperatures, so as to provide a desired resulting temperature of the combined mixture.

The stripping gas may suitably be hydrogen which is well suited for the desired stripping function and which can readily be recycled from the gas phase product of the stripping and washing step. Of course, other sources of hydrogen or other stripping gas could be used if desired.

The washing feed may suitably be Diesel, hydrotreated naphtha, LVGO or any other suitable washing substance, which could advantageously be provided from storage, from VGO liquid fractions separation (VGO), or from other treatment units such as DC, FCC, distillation, low pressure HDS units and other units or processes. In this regard, any of these sources could be regarded as external feed sources.

In accordance with the invention, the reaction feed, stripping gas and washing feed are preferably each fed to the stripping and washing zone in amounts sufficient to provide the desired separation of gas and liquid phases. Stripping gas may suitably be fed to the stripping and washing zone in an amount between about 10 and about 100 ft

3

of gas per barrel of reaction feed. Washing feed may advantageously be fed in an amount between about 5% v/v and about 25% v/v with respect to the reaction feed.

It is particularly advantageous that the gas phase produced from the separating and washing step is produced at a pressure which is within about 80 psig and more preferably is within about 50 psig of the pressure of the upstream or VR hydrocracking reaction zone, and is further therefore still at a pressure sufficiently elevated that desirable second reactions such as hydrotreatment-hydrocracking and the like can be carried out without needing to feed the gas phase to a compressor.

In accordance with the hydrocracking/hydrotreating-hydrocracking embodiment of the present invention, the gas phase from the hot separator-stripper-washing zone is fed to a second reactor for carrying out hydrotreating-hydrocracking so as to improve the distillate fraction yields and quality. The product of the hydrotreating-hydrocracking reaction step includes a kerosene or jet fuel having a smoke point between 20 and 28, which is substantially increased, by between 5 to 8 numbers in comparison with the first reaction stage product, a Diesel fraction having a cetane number between 45 and 60, which is substantially increased, by at least about 13 numbers in comparison with first reaction stage or imported fraction, and jet fuel and Diesel with a sulfur content less than or equal to about 20 ppm. The gasoline fraction is provided having a sulfur content of less than or equal to about 5 ppm.

Additional liquid product fractions from the separation-stripping-washing zone can advantageously be fractionated to recover VGO for further recycling and processing, and the liquid phase contains remaining catalyst and vacuum residue which advantageously allows the second reaction zone to be a trickle bed hydrogenating-hydrocracking reactor preferably containing effective amounts of a suitable catalyst, preferably a sulfur-nitrogen resistant catalyst selective toward aromatic saturation and alkylparaffin forming reactions. Of course, the second reaction may be any desirable high pressure reaction, and the catalyst should be selected having activity toward the desired reaction.

Turning now to

FIG. 2

, a process in accordance with the present invention is schematically illustrated.

FIG. 2

shows a first reactor system

24

for carrying out an initial hydrocracking reaction, second and third reactors

26

and

28

for carrying out a hydrotreating-hydrocracking reaction, and a high-pressure hot separation-stripping-washing unit

30

connected between reactor

24

and reactors

26

,

28

for advantageously separating the product of reactor

24

into a high pressure gas phase for treatment in reactors

26

,

28

according to the invention, and a liquid phase for further fractionation and VGO recycle.

As shown, the process advantageously begins through providing a vacuum residue feed

32

which can be fed to a heater if desired and which is then fed to first reactor

24

. The converted products from first reactor

24

are conveyed through various stages and then as feed to an inlet to hot separation-stripping-washing unit (HSSW)

30

, along with additional VGO and Diesel from an external source, hydrotreated naphtha and a feed of hydrogen as stripping gas. This combination of components forms the feed blend to unit

30

. Unit

30

produces a gas phase

34

containing, ideally, hydrogen, naphtha and Diesel fractions as well as LVGO, C

1

-C

4

hydrocarbons, H

2

S and NH

3

. The gas phase

34

, or a portion thereof, is then fed directly to second and third reactors

26

,

28

where VGO, kerosene, Diesel and naphtha fractions are subjected to hydrotreating-hydrocracking so as to increase the distillate yields and quality such us smoke point, cetane number and sulfur as desired. Product

36

from second and third reactors

20

,

21

can then be separated into gasoline, jet fuel, Diesel, unconverted VGO that is recycled, and other fractions.

A portion of unconverted vacuum residue may be separated off as fuel for heater(s), if desired, so as to provide for desired heating of the vacuum residue feed. Of course, other heating mechanisms and methods could also be used.

In addition, hydrogen is in this embodiment separated from the gas phase of product of second-third reactors

26

,

28

, preferably downstream of reactors

26

,

28

, and is purged and recycled for mixing with vacuum residue to form the

00-152

feed blend for the hydrocracking reactor system

24

, and could also be recycled as stripping gas.

The H

2

S and the NH

3

portions of gas phase

34

can be separated prior to feed to reactors

26

,

28

, if desired.

A particular advantage of the present invention is that hydrocracking reactor system

24

(R

1

), hydrotreating-hydrocracking reactors

26

,

28

(R

2

, R

3

) and stripping/washing unit

30

are all operated at substantially the same pressure such that no additional compressor equipment is required along the process stream from first reactor system

24

through unit

30

to second and third reactors

26

,

28

. Thus, equipment and other overhead costs in connection with the process of the present invention are significantly reduced while end products are advantageously low in sulfur content while nevertheless including jet fuel and Diesel fractions possessing increased smoke point and cetane number respectively.

Referring now back to

FIG. 1

, the hot separation-stripping-washing steps of the present invention are further discussed. Input to these steps includes external Diesel or VGO mixture as a washing feed, a converted Diesel fraction

12

from a first reactor (VRHCK) as a reaction feed, a liquid hydrotreated naphtha phase

38

and makeup hydrogen

16

as stripping gas. Also as shown and discussed above, HSSW zone

10

may have two zones

14

,

15

, and the gas phase

40

from zone

32

may be fed to zone

15

with gas phase

21

serving as feed to a hydrotreatment-hydrocracking or other process such as reactors of

FIG. 2

or the like. The product stream from HSSW zone

10

also includes a liquid phase

22

containing stripped VGO, vacuum residue and catalyst, as well as other liquid products which are preferably conveyed to a fractionator or vacuum distillation to recover VGO and other distillable products. The unconverted vacuum residue plus catalyst is ideal for processing at a delayed coker and the like.

The operating conditions for the hydrocracking (VRHCK) reactor system and subsequent hydrotreating-hydrocracking reactors are advantageously selected so as to maintain and utilize the pressure from the first reactor system (VRHCK) in the subsequent reactors and thereby enhance efficiency and avoid the need for additional compressor equipment therebetween. The process operating conditions for the initial reactor may be selected based upon the characteristics of the feed, for example, and these operating conditions can then be determinative of the operating conditions in the subsequent reactors. Table 1 set forth below provides examples of typical operating conditions for a VRHCK reactor (R

1

) and subsequent hydrotreatment reactors (R

2

) and (R

3

) for start of run (SOR) and end of run (EOR).

TABLE 1

R1

R2-R3

Condition

SOR

EOR

SOR

EOR

Pressure psig

3000/2900

3000/2900

2850/2800

2850/2750

inlet-outlet

LHSV h-1

1-0.7

1-0.7

0.75-1.5

0.75-1.5

Temperature o

400-450

400-450

380-350

360-380

range

Beds with

2-3

2-3

2-3

2-3

Quench

H2 partial pres.

2000-2400

2000-2400

1700-2000

1650-1950

Psig

An example of typical feed for the Hydrocracking reactor system for the process of the present invention is set forth in the last column below in Table 2. An example of typical feeds to hydrotreating and hydrocracking reactors are set forth in Table 2 columns 2 to 6.

TABLE 2

Component of the Feed to R2/R3--

Feed R1

HCN

HCGO

AGO

LVGO

HVGO

VR

API GRAVITY

52.4

20.8

23

20.2

16.5

3

NITROGEN,

280

4433

541

846

1513

0.4

wppm

SULFUR,

1.23

3.80

2.00

2.30

2.70

4.3

wt %

CONRADSON

—

0.14

0.01

0.13

0.52

23

CARBON,

wt %

DISTILL-

TBP

TBP

TBP

TBP

TBP

TBP

ATION

IBP

163

623

570

418

588

530

5

182

634

680

495

702

570

10

200

644

705

527

748

630

30

247

688

746

608

829

50

289

744

775

671

883

70

328

809

815

733

938

90

363

887

885

816

1011

95

380

911

927

859

1046

FBP

397

937

962

928

1067

As set forth above, the feed to the VRHCK reactor system preferably includes a vacuum residue, and feed to the hydrotreating-hydrocracking reactors may typically include a blend of VGO, Diesel and other components. Table 3 below sets forth characteristics of a typical VGO feed blend, and Diesel blend for the hydrotreating-hydrocracking reactors (R

2

-R

3

) in accordance with the present invention. The R

2

and R

3

feedstock could be composed of different proportions of product coming from first reactor R

1

and from external sources of components. To provide a specific example, Table 3 groups the VGO, Diesel and kerosene fractions.

TABLE 3

Reactor stages

R2/R3

R2/R3

R2/R3

INLET

VGO blend

Diesel

Kerosene

blend

API GRAVITY

16-22

28-33

27-34

SULFUR, wt %

1.0-3

0.02-2

0.01-0.1

NITROGEN, wppm

3000-15000

200-1500

0.001-0.01

CONRADSON CARBON,

0.1-0.5

—

—

wt %

BROMINE NUMBER,

4-20

0.1-20

0.1-3

cg/g

METALS CONTENT

0.01-4

—

—

(Ni + V) wppm

CETANE NUMBER

—

30-40

30-40

AROMATICS CONTENT

3-50

20-75

10-18

wt %

Smoke point

—

—

15-17

As shown, the typical reactor feed to the hydrotreating-hydrocracking reactors will have an unacceptably high sulfur content, the Diesel blend will have a cetane number of between about 30 and about 40, which is unacceptable for incorporating into the Diesel pool, and the kerosene blend will have a smoke point between about 15 and about 17, which is unacceptable for incorporating in the jet fuel pool. These fractions are therefore “out of spec” and need to be improved.

Table 4 below sets forth characteristics of a typical unconverted VGO fraction, a typical jet fuel fraction, and a Diesel fraction which are products of the hydrotreating-hydrocracking reactors.

TABLE 4

Reactor stages

R2/R3

R2/R3

R2/R3

OUTLET

VGO

Diesel

Kerosene

API GRAVITY

19-24

30-35

33-36

SULFUR, wt %

0.002-0.02

0.0001

0.00002

NITROGEN, wppm

10-20

1-10

1-2

CONRADSON CARBON,

0.01-0.05

—

—

wt %

BROMINE NUMBER, cg/g

˜0

˜0

0

METALS CONTENT

˜0

—

—

(Ni + V) wppm

CETANE NUMBER

—

45-58

48-53

AROMATICS CONTENT

3-30

20-45

wt %

Smoke point

—

—

22-25

The final process product includes VGO which is useful for FCC or lube, Diesel for the Diesel pool, kerosene for jet fuel, and naphtha for reforming units. The fractions advantageously have significantly reduced sulfur content, and the Diesel fractions have cetane number which has been increased substantially thereby making the Diesel fraction acceptable for incorporation into the Diesel pool. The kerosene fraction has smoke point and cetane number which have been increased substantially thereby making the kerosene fraction acceptable for incorporation into the jet fuel pool or the Diesel pool.

In light of the foregoing, it should be appreciated that a process has been provided for advantageously treating vacuum residue, VGO and Diesel and naphtha feed so as to sequentially remove sulfur from all the streams and increase the cetane number of Diesel fractions and smoke point in the kerosene in a process which is efficient in terms of both energy and equipment. The process therefore provides for converting readily available feeds into valuable end products.

Turning now to

FIGS. 3

,

4

and

5

, several additional embodiments of the stripping and washing (HSSW) zone

10

of the present invention, previously shown in

FIG. 1

, are described.

FIG. 3

shows a hot separation-stripping-washing unit

30

in accordance with the present invention receiving a reaction feed

12

from a hydrocracking process (R

1

). The reaction feed includes hydrogen, naphtha, kerosene, Diesel, LVGO, HVGO, vacuum residue, C

1

-C

4

hydrocarbons, hydrogen sulfide, ammonium, catalysts and contaminants. Reaction feed

12

is introduced into unit

30

, typically at an intermediate vertical position such that stripping gas

42

can be introduced vertically lower than reaction feed

12

, and washing feed

44

is introduced at a vertically higher position than reaction feed

12

. Counter-current flow occurs within unit

30

, with stripping gas

42

proceeding upwardly through the unit and washing feed

44

flowing downwardly, each performing the desired function so as to assist in producing the desired separated gas phase

34

including hydrogen, naphtha, Diesel, LVGO,HVGO, C

1

-C

4

, H

2

S and NH

3

. Also produced is liquid portion or phase

46

containing small amounts of LVGO, HVGO and particularly vacuum residue plus catalyst, that can advantageously be passed as feed to a vacuum distillation column to produce VGO to be recycled, and a vacuum residue and catalyst fraction that can be advantageously fed to a delayed coker unit or the like for further processing.

Still referring to

FIG. 3

, stripping and washing unit

30

in this embodiment is provided as two units

48

,

50

, with reaction feed

12

introduced into unit

48

. Stripping gas

42

in this embodiment is fed to a lower portion of unit

48

, and washing feed

44

is introduced to an upper portion of unit

48

. Additional stripping gas

52

is also fed to a lower portion of unit

50

, and units

48

,

50

each produce a gas phase fraction

54

,

56

which is combined to form gas phase

34

. Further, liquid

58

exiting upstream unit

48

is introduced to downstream unit

50

and, after further stripping results in liquid phase

46

identified above, suitable as feed to a vacuum distillation column

60

for separating different fractions for further treatment.

FIG. 4

also illustrates another alternative embodiment of HSSW zone

10

according to the invention. As shown, reaction feed

12

is fed to unit

30

which in this embodiment is, like in other embodiments, provided in two units

48

,

50

.

Washing feed

44

is introduced to unit

50

as shown, and stripping gas

42

in this embodiment is introduced to a lower portion of upstream unit

48

. Unit

48

produces a gas phase fraction

54

including the desired components as discussed above that are sent to unit

50

for further separation, and a liquid phase

58

which is fed to vacuum fractionation

62

. Unit

50

produces an additional liquid phase

64

which is combined with liquid phase

58

and sent to vacuum fractionation

62

, and gas phase

34

for further hydrotreatment-hydrocracking. The downstream reaction could for example be a hydrotreating reaction or a second separator zone plus a hydrotreating reaction, and in such an embodiment, additional naphtha/Diesel may be mixed with gas phase

34

to produce the desired hydrotreating reaction feed. Washing feed

44

can be from external source as shown or may be recycled from vacuum fractionation

62

if desired.

FIG. 4

shows recycled VGO

68

being mixed with external VGO, naphtha and/or Diesel to form an additional component of feed to the downstream reactors. Diesel could also be used for this purpose.

FIG. 5

illustrates a further embodiment of the process of the present invention wherein a downstream hot separator

70

is used to further treat gas phase

34

from units

48

,

50

. Hot separator

70

produces a gas phase

72

which can be fed to a hydrocracking reactor. Gas phase

72

from hot separator

70

would include hydrogen, hydrogen sulfide, ammonia, C

1

-C

4

, naphtha, kerosene, Diesel and LVGO. Liquid phase

74

is composed mainly of LVGO and HVGO. Hot separator

70

serves to further enhance the separation of phases while still maintaining the desired temperature and pressure through to the downstream hydrotreatment and hydrocracking reactors. External VGO and/or external Diesel

76

in this embodiment may be useful to mix with gas phase

34

or gas phase

72

in order to control reaction temperature in downstream reactor. External VGO and/or Diesel may also serve as washing feed, either alone or with recycled VGO as shown in FIG.

5

.

In accordance with another embodiment of the present invention, the stripping and washing zone and steps as described above can advantageously be used to integrate one or more hydroprocessing stages or reactors into other high-pressure and high-temperature circuits including but not limited to residue hydrocracking circuits and the like, thereby obtaining further benefits from the already established high-pressure and high-temperature conditions and utilizing one sequential treatment without cold separating or conventional hot separating. This is particularly advantageous as compared to conventional systems that treat products of a first reaction to obtain an intermediate product, and then separately treat this intermediate product to obtain final product.

FIG. 6

illustrates a residue hydrocracking scheme incorporating an HSSW zone

10

similar to the embodiment of FIG.

5

.

FIG. 6

shows HSSW zone

10

, a hot separator

70

, a gas phase reactor

78

, and hydrotreating and hydrocracking reactors

80

,

82

.

In this embodiment, a sequential process is established utilizing vacuum residue hydrocracking (R

1

) followed by hot separation stripping and washing

10

, conventional hot separation

70

, and gas phase hydrotreatment-hydrocracking reactors

78

(R

2

),

80

(R

3

) and

82

(R

4

), so as to produce desired products.

FIG. 6

shows a vacuum residue feed to a VR hydrocracking reactor R

1

which produces reaction feed

12

to HSSW zone

10

. Liquid phase

46

from zone

10

is fed to fractionator

62

for separating a VGO fraction

68

for recycling as washing feed to zone

10

. Hydrogen stripping gas

42

is also fed to zone

10

as shown. Gas phase

34

from zone

10

is fed to hot separator

70

for further treatment at gas phase reactor (GPR)

78

so as to produce an upgraded gas product

84

, while a liquid phase

74

from hot separator

70

is fed to reactors

80

(R

3

) and

82

(R

4

) for producing an upgraded liquid product

86

which is advantageously combined with gas phase

84

and fed to a series of separators

88

for separating a gas phase

90

, a liquid product phase

92

for additional processing, for example in a fractionation or distillation unit

95

, and a final liquid phase or bottoms

94

. As shown, gas phase

90

may be treated in separator

96

to obtain recycled hydrogen

98

for use as recycled stripping gas and to satisfy any hydrogen needs at reactor R

2

, R

3

and/or R

4

. Unreacted vacuum gas oil (DVGO) is recycled from unit

95

and may be further treated in reactor

82

(R

4

) as shown so as to enhance conversion rates to desired products. This scheme is applied in Examples 1 and 2 discussed below.

In accordance with the present invention, and as further illustrated in

FIGS. 7-10

, a large number of different systems or configurations can be established for the second, third and any additional reactors or hydrotreatment reactor systems for use in treating the gas and liquid phases from HSSW zone

10

in accordance with the present invention.

FIG. 7

shows an HSSW zone

10

product gas phase

34

which is treated in a manner selected and adapted to accomplish naphtha production from the first vacuum residue hydrocracking product.

FIG. 7

shows gas phase and reaction product

34

from hot separation-stripping and washing zone

10

being fed to a first hydrotreatment reactor

100

. First hydrotreatment reactor

100

is preferably a trickle bed reactor, and inlet

102

to reactor

100

may advantageously be a liquid-gas contactor inlet adapted to provide for trickle bed operation as desired. Further, first hydrotreatment reactor

100

is provided with catalyst, preferably several different catalysts in different beds, for obtaining a desired end result. The product stream

104

from first hydrotreatment reactor

100

may suitably be fed to a second hydrotreatment reactor

106

.

Second hydrotreatment reactor

106

in this embodiment is also preferably provided with several catalyst beds and adapted to provide a final product stream

108

which may advantageously include hydrogen, naphtha, jet fuel, Diesel fuel, and unconverted vacuum gas oil fractions. The catalyst beds for reactors

100

and

106

are labeled B

1

, B

2

, B

3

and B

4

, and these beds may advantageously contain the following types of catalysts for maximizing naphtha production: B

1

—metal trapping and coke resistant catalyst; B

2

—HDS-HDN-HAD catalysts; B

3

—KCK catalyst; and B

4

—deep naphtha desulfurization catalyst. As will be further discussed below, the catalyst beds for such a system can advantageously be selected so as to provide the highest possible conversion of vacuum gas oil to desirable naphtha product as in this embodiment, or to other desirable products.

In the system shown in

FIG. 7

, the product stream from second hydrotreatment-hydrocracking reactor

106

is advantageously fed to a cold separator (not shown) wherein unconverted gas oil (UCVGO or DVGO) is recycled for further treatment. In the embodiment of

FIG. 7

, unconverted gas oil is recycled and mixed with the intermediate product stream from first hydrotreatment reactor

100

for further treatment in second hydrotreatment reactor

106

, thereby providing for higher total conversion rates of the initial gas oil feed.

In the embodiment of

FIG. 7

, wherein maximized naphtha production is desired, it may be desirable in accordance with the present invention to provide first hydrotreatment reactor

100

and second hydrotreatment reactor

106

each with two distinct catalyst bed arrangements as disclosed.

It should be appreciated that the catalyst system as illustrated in FIG.

7

and described above is tailored to tolerate some carryover of metallic or asphaltene-type material, and also to desulfurize and hydrogenate distillates so as to provide the desired product. One example of suitable types of catalyst for use in this type of process includes those catalysts described in U.S. Pat. No. 4,520,128, which catalyst is active to help reduce the affects of asphaltene carryover. The second hydrotreatment reactor

106

in this embodiment has also been tailored to be sulfur and nitrogen tolerant and the hydrocracking catalyst for us in first bed B

3

may for example be catalyst as described in U.S. Pat. No. 5,558,766.

A system as illustrated in

FIG. 7

can be operated to provide for 80% or more conversion (by volume) of the combined vacuum gas oil from the HSSW zone

10

and external VGO, and conversion is primarily to naphtha products. Such a process would consume a large amount of hydrogen, but would also produce an extremely high quality product.

FIG. 8

illustrates an alternative embodiment of the integrated process scheme of the present invention, wherein first and second hydrotreatment reactors

110

,

112

are provided and operate or tailored to obtain a high quality Diesel oil final product. In this case and HSSW zone

10

and two hot separator

114

,

116

are used.

In the embodiment of

FIG. 8

, first reactor

110

has a first catalytic bed B

1

and a second catalytic bed B

2

, and second hydrotreatment reactor

112

also has a first catalytic bed B

3

and a second catalytic bed B

4

. In this embodiment, additional hot separators

114

,

116

are provided, each as described in connection with the stripping and washing zones and embodiment of

FIG. 5

described above, for further enhancing the desired outcome of this process.

As shown, first reactor

110

accepts as a reaction feed a portion of external VGO

118

and the liquid fraction

120

from hot separator or

114

. First reactor

110

produces an intermediate product stream

122

which is advantageously fed to second hot separator

116

as shown.

First and second additional hot separators

114

,

116

each also produce a gas phase as described above, and these gas phases are advantageously combined with an external Diesel feed

124

and fed to second hydrotreatment reactor

112

. These gas phases, as described above, advantageously contain hydrogen, naphtha, Diesel and LVGO fractions, and hydrogen sulfide and ammonia contaminants. The final product stream from second hydrotreatment reactor

112

, along with the liquid phase from second additional hot separation system

116

, may suitably then be fed to a fractionation unit and the like for obtaining the desired separate product fractions. Hydrogen obtained from this step may advantageously be recycled to various points in the process as illustrated in FIG.

8

.

It should be appreciated that the embodiment of

FIG. 8

is dedicated or tailored for the production of high quality Diesel fuel with moderate VGO conversion rates (MHCK). In this scheme, two additional hot separator systems are use, and provide for additional benefits in the first and second hydrotreatment reactor

110

,

112

.

As with the embodiment of

FIG. 7

, the catalytic beds provided in first and second hydrotreatment reactors

110

,

112

are preferably provided having specific particular functions and catalysts.

First hydrotreatment reactor

110

preferably has first catalytic bed B

1

provided with a similar metal trapping and coke resistant catalyst as the embodiment of

FIG. 7

described above. Second bed B

2

of first hydrotreatment reactor

110

may suitably have a mild hydrocracking catalyst, typically having a 50% VGO conversion rate, and this catalyst is preferably sulfur and nitrogen resistant.

Second hydrotreatment reactor

112

is also preferably provided as described above with first and second catalytic beds B

3

, B

4

, and first catalytic bed B

3

may advantageously be provided having a catalyst active toward hydrodearomatization as well as sulfur and nitrogen removal, and second catalytic bed B

4

may advantageously be a catalyst selected from naphtha hydrodesulfurization. Diesel and naphtha produced in first hydrotreatment reactor

112

are then improved in second hydrotreatment reactor

114

and conversion of VGO for such a process would typically be approximately 50% volume in first hydrotreatment reactor

112

, and between about 20 and about 30% volume in second hydrotreatment reactor

114

.

Turning now to

FIG. 9

, still another alternative embodiment of the integrated hydroprocessing scheme of the present invention is illustrated. In this embodiment, three hydrotreatment reactors

126

,

128

,

130

are provided and tailored to hydrocract for conversion to jet fuel and Diesel fuel fractions.

In this embodiment, the product stream

34

from and HSSW zone

10

as described above is fed to first reactor

126

and results in first intermediate product stream

132

. First intermediate product stream

132

may advantageously be mixed with an external or recycled vacuum gas oil feed (VGO), and this mixture can be fed to a hot separation system

134

as described above. Liquid phase

136

from hot separation system

134

can advantageously be fed to second reactor

128

, while the gas phase

138

from additional hot separation system

134

can advantageously be mixed with external Diesel and utilized as a portion of feed to third reactor

130

. A second intermediate product stream

140

results from second hydrotreatment reactor

128

, and this second intermediate product stream may suitably be fed to a second hot separation system

142

where a gas phase is produced and mixed with the gas phase of first additional hot separation system

134

and external Diesel for input to third hydrotreatment reactor

130

as desired.

The liquid phase

144

from second hot separation system

142

may be set aside for feed to a fractionation unit or the like, preferably along with liquid phase product

146

from third reactor

130

as desired.

It should be also noted that hydrogen is obtained from the final product, for example at the fractionation unit, and in this embodiment this hydrogen is preferably recycled as feed to second hydrotreatment reactor

128

as shown.

The scheme as illustrated in

FIG. 9

is ideally used for deeply hydrocracking VGO into jet and Diesel fuel, and can provide for excellent results with or without recycling of unconverted VGO. The preferred catalytic beds of reactors

126

,

128

and

130

include a metal trapping catalyst B

1

followed by a sulfur and nitrogen resistant mild hydrocracking catalyst B

2

in first reactor

126

, hydrocracking catalyst B

3

which can be the same or different in second hydrotreatment reactor

128

, and sulfur and nitrogen resistant catalyst for improving Diesel fuel B

4

and desulfurizing naphtha B

5

can advantageously be provided in the beds of third hydrotreatment reactor

130

.

In this embodiment, first and second hydrotreatment reactors

126

,

128

both include hydrocracking stages, while the third hydrotreatment reactor

130

is dedicated to hydrofinishing of jet and Diesel fuel.

Turning now to

FIG. 10

, still another alternative embodiment in accordance with the integrated processing scheme of the present invention is illustrated.

In this embodiment, products

34

from HSSW zone

10

are fed to an additional hot separation system

148

, wherein a liquid phase

150

is obtained for feeding to a hydrotreatment reactor

152

, and a gas phase

154

is separated. The liquid phase fed to reactor

152

can suitably be contacted with recycled hydrogen gas, and the product

156

of reactor

152

is mixed with gas phase

154

from hot separator

148

and sent to an additional hot separation system

158

so as to provide a gas phase

160

composed hydrogen, naphtha, kerosene, Diesel, and light VGO that are treated in a finishing reactor

162

. The liquid

164

from second hot separator

158

is mixed with product

166

of reactor

162

and sent to a separation stage to recover medium-high quality products or syncrude as desired.

In accordance with this alternative embodiment, catalysts are selected for obtaining a medium-high quality syncrude production integrating a HSSW system and a mild-hydrocracking reaction system in a vacuum residue high pressure high temperature loop. Specifically, reactor

152

has a first bed B

1

of metal trapping coke resistant catalyst, and a second bed B

2

of HDS-HDN MHCK catalyst, while reactor

162

has a finishing catalyst B

3

such as HDS-HDN (Diesel and naphtha finishing) catalyst. In this process, VGO is also converted into distillates. The resulting product quality is excellent, although somewhat lower than quality obtained utilizing other embodiments of the invention.

The systems as illustrated in

FIGS. 6-10

above are various embodiments which advantageously integrate one hydrotreatment-hydrocracking process into residue hydrocracking circuit utilizing the hot separation stripping and washing (HSSW) method of the present invention. These different processes each advantageously provide for savings in economy through utilization of already existing high pressure and high temperature conditions, and through treating various fractions in a single reactor, thereby providing for savings in equipment as well.

It should be readily appreciated that

FIGS. 3

,

4

and

5

illustrate variations of the stripping and washing steps which are all well within the broad scope of the present invention, and which all advantageously provide for high temperature and high pressure separation of a reaction feed into a gas phase and liquid phase containing the desired components for subsequent processing in one, two or more stages of hydrotreatment-hydrocracking, and that

FIGS. 7-10

illustrate variations of sequentially integrated hydrotreatment-hydrocracking stages using the hot separation-stripping and washing systems of

FIGS. 1

,

3

,

4

and/or

5

or variations thereof.

EXAMPLE 1

In order to illustrate the advantageous results obtained in accordance with the present invention, two processes were run. The first process used a vacuum residue hydrocracking reaction zone followed by hot separation-washing, and a sequentially hydrotreat-hydrocrack of the gas phase produced by the HSSW system. Liquid phase from the HSSW system was stored for further use. In this first process, naphtha, Diesel and VGO fractions were produced and then treated in second and third reaction stages to upgrade yields and quality. This scheme was call SEHP

1

.

The second process used the same vacuum residue hydrocracking stage (R

1

) (same operating severity). The product was fed to an HSSW system; the liquid product from the bottom of the HSSW system was stored, and the gas phase produced from the top of the HSSW zone was fed to a second separation system at the same pressure and at 40° C. less temperature, as shown in FIG.

5

.

Gas phase from the second high pressure hot separation system (HSSW) containing hydrogen, naphtha, Diesel, LVGO, C

1

-C

4

hydrocarbon, H

2

S and NH

3

, was fed to a hydrotreating zone (R

2

). The pressure of the gas phase was within about 80 psig of the pressure of the residue hydrocracking reactor outlet (R

1

); the liquid phase from the second hot separation system (bottom) contained small amounts of Diesel, LVGO and HVGO and was blended with recycled hydrogen and fed to a hydrocracking reactor zone (R

3

) operated within about 80 psig of the pressure of the residue hydrocracking reactor outlet (R

1

). The products from the hydrotreating reactor and hydrocracking reactor were combined and fed to a fractionation tower. This scheme is called SEHP

2

.

Table 5 sets forth the results of these two process schemes using the same temperature space velocity and reactor volume, but using one or two stage of high pressure and high temperature separation. In the second case (two high pressure separators) a high purity of hydrogen makes up gas if needed, are recycled to the hydrocracking reactors. Table 5 shows that yield and quality are improved from the yield and quality fed to the HSSW system. Using one or two stages of separation improve the yield and quality of naphtha, kerosene and Diesel.

TABLE 5

Without

With

SEHP

SEHP1

with SEPH2

Naphtha yields wt %

10

18

15

Jet Fuel wt %

5

13

12

Jet Fuel smoke point

16

22

24

Diesel yields wt %

13

20

29

Diesel Cetane Number

44

55

58

VGO yields

56

28

24

650° F.+ Conversion (R2/R3)

0

50

(+R4)

56

950° F.+ Conversion (R1)

90

90

90

SEPH

1

HSSW (

FIG. 2

without External Feedstock)

SEPH

2

HSSW+HSS (

FIG. 6

without External Feedstock)

As shown, the process conducted with high temperature and high pressure stripping and washing followed by a simple hydrotreating and hydrocracking stage (SEPH

1

) produced more C

1

-C

4

/naphtha and VGO and less kerosene-Diesel products than the scheme with HSSW followed by a hot separation system (SEPH

2

). In addition the product quality is better in the latter process. These processes in accordance with the present invention produced an increase in smoke point of 6-8 numbers, and produced an increase in cetane of 11-14 number (second and third columns) in comparison to the vacuum residue hydrocracking product without integrating the SEHP technology (first column).

EXAMPLE 2

In order to further illustrate the advantageous results obtained in accordance with the present invention, two modes of application HSSW with hydrotreating and hydrocracking processes were run with the same initial vacuum residue hydrocracking reaction stage, but with different ways of hydrotreating the reaction feedstock and the external feedstock. SEHP

3

designates one mode where the gas phase produced in the stripping-washing separation stage is blended with 20% vol. external VGO, 15% vol. external Diesel and 10% vol. naphtha fractions, relative to the VGO Diesel and naphtha leaving the hydrocracking reactor and entering the HSSW system. Then the gas phase is slightly cooled (but at the same pressure) by blending outside feeds, and the blend, is sent to hydrotreating-hydrocracking reactors (R

2

/R

3

). In the second process or mode (called SEHP

4

), the same vacuum residue hydrocracking severity is used in the initial reactor, otherwise the same amount of naphtha, kerosene, Diesel and VGO are sent to the HSSW system as in the previous case (SEHP

3

) and the same amount of external CGO Diesel and naphtha is added but in different way.

The gas phase produced in the HSSW system is partially cooled by adding 20% vol. external VGO (same pressure) sent to a second high pressure, hot separation system. From there a gas phase is separated from the top and blended with the 15% vol. external Diesel and 10% external naphtha. That gas phase stream, rich in hydrogen, naphtha, kerosene, Diesel and light VGO, is sent to a hydrotreating reactor. The liquid phase obtained from the second hot separation system is blended with recycled hydrogen and sent to hydrotreating-hydrocracking reactors (R

2

/R

3

) operating at substantially the same pressure as the previous reactor (R

4

). In addition, the volume of all the reactors (R

2

/R

3

/R

4

) and the temperature are the same as the previous case (SEHP

3

). Table 6 sets forth the results of this process, identifying the SEHP

3

process with one hydrotreating-hydrocracking stage and “SEHP

4

” as the two stage hydrotreating-hydrocracking process. Notice that both schemes use the same feedstock and the same operating conditions and the same separation-stripping and washing stages.

TABLE 6

Blend*

SEHP3

SEHP4

Naphtha sulfur wt ppm

1000

30

2

Kerosene smoke point

16

22

24

Diesel CI

41

47

52

VGO Sulfur wt %

3.41

700

400

VGO conversion wt %

0

82

87

SEHP3 according to FIG.

2

SEHP4 according to FIG.

6

*The blend is produced by mixing the vacuum residue hydrocracking products plus the external VGO, Diesel and naphtha.

Most of the blend properties are better than the products from vacuum residue conversion (R

1

outlet). By sequentially integrating the HSSW plus the hydrocracking reactor, the VGO conversion increases and the quality of the unconverted VGO is improved. The same is true for naphtha and Diesel as shown in Table 6.

The process conducted with high pressure stripping and washing and one hydrotreating stage accomplished and important reduction in sulfur content and improved Diesel, cetane number and kerosene smoke point. However, the process carried out utilizing two hydrotreating stages resulted in a greater sulfur reduction and higher cetane number and smoke point. The SEHP

4

mode also experienced a greater conversion rate for the VGO fraction.

EXAMPLE 3

This example illustrates washing and stripping separation (HSSW) in accordance with the present invention.

FIG. 11

shows a mass and energy balance of an HSSW zone or vessel according to the invention, showing flowrate S (tons/hour) and temperature for reaction feed, stripping gas, washing feed, gas-phase product and liquid phase product. This shows beneficial separation without cooling, or loss of pressure.

Top and bottom streams were analyzed from this process utilizing 50 ton/hour of hydrogen stripping (approximately 10% wt.) and using 209 ton/hour of washing feed (approximately 15% vol.) and using for comparison a hot separation process with no stripping or washing. Table 7 sets forth the results.

TABLE 7

Inlet

Top

Bottom

Top*

Wash Oil

HS System/Streams

23

303

45

x

H

2

S

193

200

1

160

0

NH

3

26

26

0

14

0

H

2

O

46

46

0

34

0

C1/C2

348

378

3

345

0

C3/C4

236

234

2

212

0

Naphtha C5-200 C.

884

878

6

856

0

Mid Dist 200-350 C.

1638

1686

95

1334

143

VGO 350-500 C.

1327

1131

362

967

166

residue 500 C.+

273

12

261

106

0

solids

114

0

114

0.1

0

Top* is the gas phase using just a hot separation

Table 7 shows hydrogen stripping and washing resulted in more gas phase production. The H

2

S, ammonia, C

1

-C

4

, naphtha and Diesel are stripped from the bottom (liquid) to the top (gas phase), and the vacuum residue and the solid are washed out from the gas phase. The total amount of products leaving the HSSW system include the amount of Diesel and VGO used for washing and the hydrogen employed for stripping. Comparing the column “Top” and “Top*”, the difference is highlighted as between the HSSW system of the present invention and conventional hot separator. The results prove that the process of the present invention HSSW reduces the vacuum residue and catalyst carryover into the gas phase, and reduces the H

2

S, ammonia, C

1

-C

4

and the naphtha going into the bottom or liquid phase to the vacuum fractionation tower. This reduces the cost of the latter. The Diesel and VGO are advantageously fractionated in the vacuum tower and partially pumped back to the stripping zone in the HSSW system in a closed loop.

EXAMPLE 4

This example demonstrates the difference between separate or stand-alone processes and the integrated processes of the present invention. A conventional scheme was established, and is illustrated in

FIG. 12

, wherein a first vacuum-residue hydrocracking, with distillation step, is carried out to obtain VGO which, along with external VGO, is treated in a stand alone hydrocracking process (reactors and all other units). Diesel from the initial vacuum residue hydrocracking step is mixed with other Diesel for hydrotreatment. The properties of products from the initial hydrocracking stages for this scheme are illustrated in Table 8 below under the heading “Scheme II”.

TABLE 8

Scheme I

Scheme II

Feed Components, BPD

Naphtha from VRHCK

15,142

0

Middle Distillates from VRHCK

23,476

0

VGO from VRHCK

17,864

17,864

Straight run HVGO

40,940

40,940

Total Feed Throughput

97,422

58,804

Fuel properties

API Gravity

23.9

14.3

Sulfur, wt %

2.12

2.78

Nitrogen, ppm

4060

3542

Conradson Carbon, wt %

0.25

0.56

Metals (Ni + V), ppm

1.5

3.6

TBP Distillation, C.

10 vol. %

157

375

30 vol. %

286

436

50 vol. %

389

478

70 vol. %

445

519

90 vol. %

516

604

Total Pressure R

1

Inlet, psig

2,800

2,400

H

2

Pressure R

1

Inlet, psig

2,050

2,200

Total Recycle Gas Rate, scfb

15,030

20,000

LHSV HDT Reactors, v/v/h

0.9

0.9

LHSV HCK Reactors, v/v/h

0.33

0.98

Cycle Length, months

24

24

A scheme was also established for comparing the process of the present invention to the conventional process. In this scheme, schematically illustrated in

FIG. 13

, all the product from vacuum residue hydrocracking plus external Diesel and VGO (same amount as in Scheme II), is sent to an HSSW system, and then the gas phase is sent to hydrotreating-hydrocracking reactors (just the reactors and nothing else) to be processed. The combined product stream from the sequentially integrated process that includes distillation steps is also set forth in Table 8 above under the heading “Scheme I”.

The Scheme II process treats all naphtha and Diesel separately from the hydrocracking process, while the Scheme I process treats all distillates in a single integrated reactor system (with a single investment in equipment). Table 9 below sets forth final results in terms of conversion, and Table 10 below sets forth final results in terms of product quality.

TABLE 9

SHP Scheme I

Scheme II

Performance

VGO Conversion, vol. %

86.7

97

Hydrogen Consumption, scfb

2177

2350

C5+ Liquid Yield, vol. %

109.7

114.8

Product Yields (based on Feed), wt. %

Hydrogen Consumption

3.62

3.66

H

2

S

2.25

2.96

NH

3

0.49

0.44

H

2

O

0.46

0

C

1

-C

2

0.61

0.72

C

3

-C

4

2.64

3.97

Light Naphtha

5.49

6.86

Heavy Naphtha

16.46

17.21

Jet Fuel

38.54

38.3

Diesel

28.68

30.03

VGO

8.02

2.7

TABLE 10

SHP Scheme I

HCK Scheme II

Light Naphtha (C5-85° C.)

API Gravity

81.5

80.5

Sulfur, ppm

0.5

0.5

Nitrogen, ppm

1

1

Octane Number, RON

c

78

78

P/N/A, vol. %

84/14/2

83/14/3

Heavy Naphtha (85-150° C.)

API Gravity

52.2

57.8

Sulfur, ppm

0.5

0.5

Nitrogen, ppm

1

1

Octane Number, RON

c

60

61

P/N/A, vol. %

40/54/6

43/49/8

Jet Fuel (150-265° C.)

API Gravity

42.5

42.8

Sulfur, ppm

5

2

Nitrogen, ppm

1

1

Aromatics, vol. %

25-21

27-23

Flash Point, F

115

120

Freeze Point, F

−63

−60

Diesel (265-350° C.)

API Gravity

37.7

35

Sulfur, ppm

10

5

Nitrogen, ppm

1

1

Cetane Number

58

50

Cetane Index

58

53

VGO (350 C.+)

API Gravity

26

26.6

Sulfur, ppm

20

10

Nitrogen, ppm

5

5

UOP K

12.8

12.6

Table 9 shows that the use of a stand alone hydrocracking unit results in a higher conversion (Scheme II 97% VGO conversion) than in the integrated system (Scheme I:86.7% VGO conversion). This is due to conversion in Scheme I process of all the product, and the results are limited by the maximum commercial vessel size available. Scheme I also provides a higher conversion to gasoline, while Scheme II provides a higher conversion to kerosene Diesel. Hydrogen consumption is greater in the stand-alone system of Scheme II.

From Table 10, it should readily be appreciated that both products have outstanding qualities. Of course, the product obtained utilizing Scheme I in accordance with the present invention is obtained along with a cost savings due to efficient utilization of high pressure and temperature conditions and further through less equipment costs.

Light naphtha properties for Scheme I and Scheme II were almost identical. In most cases, this naphtha product could be blended directly into the low octane pool. If additional octane is required, the light naphtha stream is an excellent feed to isomerization processes. Sulfur and nitrogen contents of the heavy naphtha fractions are very low in both schemes, and the naphthalene plus aromatics content is almost 60%, making these streams excellent feedstocks for naphtha reformers and the like.

The kerosene streams, Diesel streams and VGO streams from both schemes also have very desirable properties. Of course, as set forth above, the process of the present invention provides for substantial savings in operating cost and equipment investment cost as compared to the non-integrated Scheme II process.

This invention may be embodied in other forms or carried out in other ways without departing from the spirit or essential characteristics thereof. The present embodiment is therefore to be considered as in all respects illustrative and not restrictive, the scope of the invention being indicated by the appended claims, and all changes which come within the meaning and range of equivalency are intended to be embraced therein.

Number	Name	Date	Kind
3365388	Scott, Jr.	Jan 1968	A
3607734	Stafford, Sr.	Sep 1971	A
3733260	Davis et al.	May 1973	A
3806444	Crouch et al.	Apr 1974	A
4010010	Ward	Mar 1977	A
4414103	Farrell	Nov 1983	A
4422927	Kowalczyk et al.	Dec 1983	A
4822480	Harandi et al.	Apr 1989	A
4954241	Kukes et al.	Sep 1990	A
5110444	Haun et al.	May 1992	A
5110562	Haun et al.	May 1992	A
5447621	Hunter	Sep 1995	A
6106694	Kalnes et al.	Aug 2000	A
6190535	Kalnes et al.	Feb 2001	B1
6296758	Kalnes et al.	Oct 2001	B1

Process scheme for sequentially hydrotreating-hydrocracking diesel and vacuum gas oil

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