Hydro Cracking Tech For Middle Dist

PEP Review 98-7 HYDROCRACKING TECHNOLOGY FOR MIDDLE DISTILLATE By David Netzer (OCTOBER 2001)

ABSTRACT The objective of this review is the production of middle distillate, 300-650°F cut point, while maximizing diesel (550-650°F) yield. The new UOP’s new HyCycle (TM) Unicracking hydrocracking technology is the basis of the evaluation. US Patents 5,980,729, 5,885,440 and a 2001 NPRA publication [R98-07-001] have been served as a basis for evaluating the technology. The key attributes to this technology are: •

Reverse arrangement of reactors system. The hydrocracking, with low conversion per pass, about 33%, comes first followed by hydrotreating. The preheated feed enters the hydrotreating reactor along with hydrocracker outlet product, for very deep sulfur and nitrogen removal.

•

The cracked products are separated from the reactor loop at about 1,800 psig in hydrogen rich vapor phase. Product is condensed outside reactor loop prior to conventional products recovery. An overall conversion of above 98% and below 650°F cut point is achieved.

•

Reactor loop nominal pressure of about 1,800 psig as opposed to 2,400 psig in more conventional double reactor system.

•

For 21.5 API, 2.5 wt% sulfur, 980°F TBP cut POINT VGO, the yield is 38 vol% diesel, 42 vol% kerosene and 29 vol% naphtha. The equivalent yield in conventional double reactor system with 97% conversion to 650°F cut point is 19 vol% diesel, 54 vol% kerosene and 36 vol% naphtha.

•

The system comprises a finishing reactor, and naphtha product, meeting the 1 ppm sulfur specification for typical downstream processing such as catalytic reforming.

•

The hydrogen consumption for the above yield basis is reduced from 1,950 Scf/bbl in conventional double reactor to 1,700 Scf/bbl in HyCycle.

•

It is thought that the capital cost investment will be reduced by about 10-14% subject to more detailed cost comparison.

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INTRODUCTION Middle distillates, diesel and kerosene are the more important refinery products for Asian countries like China and India, and other countries with emerging economy and infrastructure. Even in Western European countries, the ratio of demand of middle distillate to gasoline products is considerably higher compared to the ratio in North America. The newly emerging HyCycle(TM) hydrocracking technology as developed by UOP is geared toward maximizing diesel yield, and this is the focus of this report. The hydrocracking process and chemistry are discussed along with the economics of hydrocracking VGO to middle distillates in PEP Report 211, Hydrocracking (1994). Hydrocracking of residual oils is discussed in PEP Report 228, Refinery Residue Upgrading (2000). The model used for this evaluation is a feed rate of 35,000 BPSD of (5,160 TPD, 5665 M3/day) of combined AGO and VGO (Atmospheric gas oil and 980°F TBP cut point vacuum gas oil). These originate from atmospheric and vacuum fractionation of Arabian Light, crude oil, 34.5 API (SG=0.852). This capacity is based on an assumed 168,000 BPSD (350 days per year, 8.0 MMTPY) grass roots refinery, which is judged to represent an average future size refinery to be built in emerging markets. The Arabian Light represents a common world benchmark of crude oil. The results of this evaluation can be adjusted to many other potential hydrocracking feeds and capacities while maintaining a reasonable accuracy. The size range of future hydrocracking units is expected to be in the range of 20,000-50,000 BPSD per single train of production. Diesel Specifications Diesel fuel regulations over the next decade are reducing the sulfur content in all the major markets to approach “sulfur free” fuel in order to reduce vehicle NOx and particulate emissions. Parts of Europe have led the sulfur reduction. For instance, since 1991 Swedish Class 1 diesel fuel has had <10 ppm sulfur, typically 1-3 ppm with cetane rating >51 [R98-07-005]. Diesel sulfur content in Europe will be reduced to 50 ppm in 2005 (European Union, Stage 4 regulation). The European commission’s (EC) latest Auto-Oil II discussion paper proposes a 10 ppm sulfur Ultra Low Sulfur Diesel fuel be phased in starting on January 1, 2007 with a 10% supply requirement [R98-07-004]. In the U.S., the Environmental Protection Agency has set a specification of 15 ppm sulfur for June 1, 2006 [R98-07-006]. This regulation, concurrent with implementation of new gasoline regulations, is being challenged in federal court by the National Petrochemical and Refiners Association [R98-07-007; R98-07-008; R98-07-009]. The U.S. specification is currently 500 ppm sulfur with aromatics <35 wt% (10 wt% in California with polyaromatics 1.4 wt%) and cetane index >40 (>48 in California) [R98-07-002]. In order to provide 15 ppm sulfur fuel at the pump, U.S. refiners may have to produce 1- ppm sulfur fuel at the refinery due to contamination in the transportation system. Japan is fine-tuning its program similar to the European and U.S. programs and tightening NOx and particulates standards for diesel powered cars and small trucks [R98-07-003]. South Korea is making major cuts in diesel sulfur levels later in the decade. The Association for Latin American Refiners has proposed diesel sulfur and cetane specifications of 2000 ppm sulfur and 47 respectively for 2005 [R98-07-003]. So far, the EC proposal and U.S. regulations have only defined the sulfur concentration; other properties such as cetane number and aromatics content, have not yet been proposed. However, the Category 3 diesel fuel specifications proposed under the automaker’s “World Wide 2 PEP REVIEW 98-7

Fuel Charter” suggest sulfur <30 ppm with cetane number >55 (or Cetane Index >52), total aromatics <15 vol% with polyaromatics <2 vol%, and a density of 820-840 kg/m3 [R98-07-005]. HyCycle Unicracking-Key Process Features The HyCycle Unicracking process is considered by many as a step change in technology for maximum diesel yield in hydrocracking. The concepts of this process are described in three patents [US 5,885,440; US 5,980,729; WO 97/38066]. This is a flow scheme for maximizing the yield of high-quality diesel fuel. The quality of the diesel fuel is measured in terms of Cetane(1) Index, aromatic content, sulfur content and nitrogen content. HyCycle cracking process equipment features include [R98-07-001]: •

Separator/finisher

•

Back-staged series flow reactors

•

Novel fractionator design

Process and catalytic features include: 1. Low per pass conversion of 20-40% with high overall conversion of over 98% 2. Hydrogen consumption is reduced up to 20% compared to conventional technology 3. Selectivity to higher boiling products is high. Up to 5 vol% more yield of middle distillate with a 15% shift towards diesel compared to other full conversion maximum diesel processes 4. Operating pressure is 25% lower than conventional gas oil hydrocracking 5. Hydrogen partial pressure is lower 6. Space velocity is higher than conventional reactors but catalysts life and product quality are unchanged. The HyCycle process operates at low per pass conversion(2) in the reactor loop, which minimizes the undesired cracking reactions, particularly of the diesel range molecules, and also minimizes coke formation reactions, which reduce catalyst activity. Selective ring opening in HyCycle shifts the equilibrium toward a more favorable regime for producing high-quality diesel fuels. The first key feature of the HyCycle Unicracking is the higher selectivity toward diesel product. This shift in selectivity avoids cracking of diesel product while at the same time increases saturation of aromatics. This is estimated to result in a net reduction of about 15% in hydrogen consumption, in this particular application, compared with the present state of the art of hydrocracking technology. A second key feature of the HyCyle process is the 25% reduced operating pressure and the higher space velocities obtained when compared with the present state of the art in hydrocracking technology. This step change technology improvement can be accomplished without compromising on catalyst life or diesel product quality. A third key feature of the HyCycle Unicracking, and assuming a FCC unit in parallel to hydrocracking, is the potential synergism in hydrotreating and aromatic saturation of light cycle oil (LCO) produced at the FCC unit. The LCO is high in aromatic low valued diesel range product that would be upgraded to high-quality diesel fuel.

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The design of the HyCycle combined with HPNA, heavy polynuclear aromatic, management concept. The HPNA is known to be one of the undesirable hydrocracking products. The HPNA separates from the cracked liquid at lower temperature and creates severe maintenance in the colder sections, resulting in excessive down time of the hydrocracking unit. Using of hot enhanced oil separation, at near reactor exit conditions, combined with downstream purge of HPNA results in increasing catalyst cycle time and overall profitability. The enhanced separator is using internal high pressure stripping with hydrogen, of diesel from liquid unconverted oil (UCO) product. The enhanced separator /stripper recovers the distillate products as vapor at reactor loop pressure. Liquid, consisting of unconverted oil from the separator, depleted of distillate, recycles to the hydrocracker reactor within the hot high pressure loop. A small portion of the unconverted oil in the “hot loop” high pressure cycle is let down from reactor loop pressure to product fractionation, operating at close to atmospheric pressure this let down avoids build up of HPNA in the reactor loop. The enhanced separator is combined with a finishing reactor. Traces of olefins that exist in hydrocracking products could recombine with H2S and form mercaptans. The finishing bed hydrotreats these mercaptans and bring down the sulfur content of the naphtha range material, to under 1 ppm. A range of catalysts are available from UOP for the HyCycle process [R98-07-001]. These are zeolite based catalysts impregnated with metals on a proprietary support material. The optimum balance between cracking and hydrogenation is judged on a case by case basis. The DHC-32 LT catalyst can be operated for maximum diesel and no jet fuel or in a flexible mode where naphtha and jet fuel are both produced along with increased diesel. The DHC-43 LT catalyst is offered for maximizing the yield of jet fuel. Maximum naphtha yield is produced with the HC-170 LT catalyst. The reactor yields of the HyCycle process, when using a flexible, mixed mode catalyst assumed in this Review, are compared with a conventional process in the table below. The conventional hydrocracker is the unit described in PEP Report 211 for middle distillates. The same Arabian Light VGO is processed in both cases. The single pass conversion of the conventional process of Report 211 is 66%. The single pass conversion of the HyCycle process in 30%. This comparison shows the yield for identical boiling range of diesel fuel is over 17% points higher for the HyCycle process than for the conventional process. Incremental diesel is produced at the expense of lighter products and fuel oil.

Product

Yield, wt% FF Conventional

HyCycle

Difference, %

Fuel Gas

0.84

0.75

-0.09

LPG

4.64

3.2

-1.44

Light Naphtha

8.22

5.70

-2.52

Heavy Naphtha

20.14

16.60

-3.54

Kerosene (Jet Fuel)

46.90

37.78

-9.12

Diesel

16.72

34.60

+17.88

Fuel Oil

2.69

2.67

-0.02

The HyCycle Unicracking process offers flexibility in staging the capital investment. As an initial operation, the HyCycle can operate in a once through, partial conversion mode, avoiding 4 PEP REVIEW 98-7

the enhanced oil separation and post reactor. This mode of operation produces unconverted low sulfur heavy fuel oil as a co-product in addition to diesel and lighter products and requires a lower capital investment. As relative demand for diesel increases with time at the expense of demand for low sulfur heavy fuel oil, the system can be modified to maximize diesel product as described above. Design Basis The HyCycle process currently produces 10 ppm sulfur, diesel product from Arabian gas oils. The diesel fuel also has a high cetane rating, low aromatics content and a density within the Category 3 range of the automaker’s “World Wide Fuel Charter” [R97-07-005]. Design assumptions •

Make up hydrogen at 99.9 vol% is available at 340 psig from battery limits.

•

Hydrocracking conversion per pass 33% and 650°F TBP cut point

•

Nominal hydrocracking and hydrotreating: temperatures 700°F pressure 1,850 psig.

•

Hydrogen at recycle compressor outlet: 90 mol%, for establishing purge rate.

•

Operating temperature of finishing reactor 625°F.

•

Sulfur content of naphtha material 1 ppm.

•

Heat of reaction 45-50 Btu/Scf of hydrogen as consumed.

•

Yield structure as shown

•

Let down of to recycle of UCO (unconverted oil) from high pressure loop to low pressure is 0.1.

The HyCycle hydrocracking plant comprises two (2) sections: Section 100, hydrogen make up compression system reaction loop, and Section 200, product recovery. Other associated process areas such as light end processing, naphtha splitter, sour water stripping, rich amine regeneration, sulfur recovery, tail gas unit and waste water treating are considered outside the battery limits (OBL) of the hydrocracking. The rationale is that light gases, naphtha, sour water, H2S, ammonia, and waste water are generated in other areas of the refinery such as crude unit and distillate hydrotreating. Therefore, all these streams are combined into a single central operation of naphtha splitting, purge hydrogen recovery, sulfur recovery and waste water treating. Make up hydrogen compression, and reactor loop, Section 100 Hydrogen, 99.9 mol% is produced from steam reforming of light naphtha (C5/C6), LPG, or natural gas from OBL. The assumed delivery pressure is 340 psig (25.0 kg/cm2-a) and containing less than 100 vol ppm of CO. This hydrogen is dedicated to the HyCycle hydrocracking. Any impure hydrogen from other sources, such as catalytic reforming will be purified prior to entering the hydrocracking reactor loop. The hydrogen is compressed in 3 stage reciprocating compressor to 2,050 psig (145.2 kg/cm2-a), with interstage air coolers to 150°F (65°C) and trim water coolers to 100°F (38°C) in stages one and two. Only partial hydrogen cooling is incorporated in to the exit of 3rd stage. Three 50% motor driven compressors are assumed, two operating and one stand by. The reactor loop encompass the key design features of the HyCycle. The ultimate conversion(3) is 98.5% and conversion per pass(2) , at 650°F (343°C) cut point, is designed for 5 PEP REVIEW 98-7

33% conversion per pass in high pressure loop. The ultimate one pass conversion of 83% in the long loop is achieved after accounting for low pressure product fractionation producing and recycling of unconverted oil. Average hydrocracking pressure, at middle of the reactors cycle, is maintained at 1,875 psig (133.0 kg/cm2-a). H2S and ammonia are removed inside the high pressure loop by a water wash of ammonia followed by the use of a lean amine, about 50 wt% MDEA solution for H2S. Loaded (rich) amine solution is routed to OBL for regeneration where it is combined with other amine streams. The catalytic portion of the reaction loop is based on information published by UOP at NPRA conferences, which was not intended for process analysis or design. The Hydrocracking (HC) reactor vessel has an assumed inlet temperature of about 700°F (371°C) which slightly increasing with time. Hydrotreating (HT) reactor vessel has an assumed inlet temperature of about 690°F (365°C) which also slightly increasing with time. In addition, there is a Finishing Reactor, which enhances the separation of distillate product vapor via hot stripping with hydrogen. It is assumed that trace sulfur removal, down to below 1 ppm for the naphtha material. The naphtha material resulted from the Finishing Reactor needs no further hydrotreating, if intended for downstream catalytic reforming of naphtha (CCR). This Finishing Reactor is combined with UOP proprietary enhanced hot separator, which operates close to pseudo critical conditions and uses hydrogen stripping of residual diesel material. The assumed separation and finishing reactor temperature is 625°F (330°C). The actual separation temperature would vary, depending on VLE (vapor liquid equilibrium) and overall heat balance. The hydrotreating and hydrocracking reactors have a design margin to work as a once through mode with 83% conversion. Under this scenario reactors will operate in a temperature range of 725-800°F and at pressure close to the design pressure. The assumed space velocities, based on cold hydrocarbon liquids is 1.15/hr in hydrocracking section and 2.3/hr at hydrotreating section. The design pressure of the reactors is set at 10% above normal operating pressure of the hydrocracking reactor. The recycle gas compressor, is a centrifugal compressor, driven by a multi-stage condensing steam turbine using steam at 600 psig/750°F, condensing to 3.5” Hg vacuum, 120°F. The turbine is equipped with an uncontrolled extraction port at 150 psig. The extraction steam option increases overall operating flexibility. Recycle gas cooling is provided by air coolers which cool the gas down to 140°F (60°C) assuming an ambient air at 100°F (38°C). A fired heater is used to provide the heat deficiency to the reactor loop. In most hydrocracking cases firing duty is provided to a mixture of hydrogen, preheated feed and unconverted oil. In this particular design situation, using hot UCO (unconverted oil) recycle and reversed reactors arrangement, only recycle hydrogen will be heated. The heater is equipped with low NOx burners and has no air preheaters. Therefore a de-NOx system is not needed to meet environmental regulations. Since the recycle heater is much smaller than the fractionation heater, excess heat in flue gas is recovered to preheat boiler feed water to be delivered for steam generation at 600 psig (42 kg/cm2-g) and 150 psig steam (10.5 kg/cm2-g) in the fractionation heater. The flue gas is released to a common stack at OBL at about 350- 400°F (176- 205°C). The firing duty will be provided by low sulfur refinery gas, which is expected to be rich in hydrogen, (probably over 50 vol%) and available at 80 psig (5.6 kg/cm2-g). Gas Oil, (AGO and VGO) feeds are assumed available at 420°F (215°C) from vacuum and crude units and after heat recovery. For base case design assume that 50% of VGO is coming from storage at 180°F and 50% from the vacuum and crude units, thus the combined temperature is about 300°F. The system, is designed for gas oil feed temperature ranging from 180°F to 420°F.

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Power recovery turbine, by reversed pumps, will be used only on one pump of the service. No power recovery is applied for spare pump and normal let down valve is used when spare pump is in service. About 300 kwh of power recovery is the assumed minimum size to economically justify power recovery reversed pump system. In general, hydrocrackers are designed for high level of heat integration in order to reduce energy cost and steam generation in hydrocracking reactor loops is un-common. In this particular case the level of heat integration is reduced while introducing steam generation to move heat load from one section of the plant to the other. This is thought to result in higher operating flexibility, at a slightly reduced thermal efficiency and probably reduced capital cost. All steam generated in the high pressure loop, will go through analytical hydrocarbons detection, through a “suspect” condensate system. Depending on specific application the “suspect condensate” can be used for water wash service for ammonium sulfide in the cold section of the hydrogen recycle. Low pressure “suspect” steam can be used for stripping services in the fractionation, and ultimately will end in waste water. Therefore this practice of steam generation avoids any potential contamination of the boiler feed water system. Product Recovery Section 200 The vapor recovery consists of mid pressure 320 psig (23.5 kg/cm2-a) V-107 and V-201 drums where hydrogen rich flash gas is recovered, and usually, depending on refinery configuration, could proceed to hydrogen recovery by PSA along with other hydrogen purge streams. The LPG stripper operates at 120 psig (9.5 kg/cm2-a) recovers C3/C4 as liquid product and releases light gas (methane and ethane) with residual propane to refinery fuel gas. Live steam at 150 psig is used for the stripping. The liquid recovery uses steam stripped fractionation, in an atmospheric column, with a nominal overhead pressure of 10 psig (1.7 kg/cm 2-a) and bottom pressure of 30 psig (3.1 kg/cm2-a). The fractionator is equipped with two pump-around loops for diesel and kerosene with 150 psig and 50 psig steam generation respectively. As an alternate, the heat released in the pump-around loops, particularly the kerosene pump-around could be used for downstream naphtha fractionation which is out of the scope of this evaluation. The diesel and kerosene products are drawn through side strippers using 50 psig steam. A HPNA purge, at a rate of about 500 BPSD, 1.5% of feed rate, is drawn from the UCO produced in the products fractionator, prior to recycling to high pressure loop. The thermal duty to the fractionation is provided by a fired heater with dual firing capability, using refinery hydrogen rich fuel gas or low sulfur fuel oil such as UCO. No air preheater is applied, nor de-NOx system. Steam at 600 psig (42 kg/cm2-g) and 150 psig (10.5 kg/cm2-g) will be generated for enhanced heat recovery, while flue gas is exhausted at 400°F (205°C). The utilities to be provided or exported to/from OBL are: •

Electric power 6,000 volt, over 200 KW service

•

Electric power 380 volt 5-200 kw service.

•

Electric power 220 volt 0-5 kw service.

•

Steam 600 psig 750°F

•

Steam 600 psig saturated, 488°F.

•

Steam 150 psig 0-100°F superheat

•

Steam 50 psig 0-100°F superheat

•

Boiler feed water, deaerated at 250°F and 750 psig 7 PEP REVIEW 98-7

•

Cooling water at 90°F and 50 psig and maximum allowed return of 115°F and minimum return pressure of 30 psig.

•

Plant air at 90 psig

•

Instrument air at 90 psig

Table 1 PRODUCTS YIELD AND QUALITY

WT%

Vol% (TBP cut points)

H2S

2.44

NH3

0.10

CH4

0.35

C2 H6

0.40

C3 H8

1.30

C4 H10

1.90

C5

2.20

3.01

C6

3.50

5.01

C7 –310°F (154°C) Heavy Naphtha

16.6

20.74

310-550°F (154-288°C) Kerosene

37.78

42.47

550-650°F (288-343°C) Diesel

34.6

38.46

650°F + (343°C +) Heavy fuel oil

1.50

1.46

102.67

111.15

Total

Light Naphtha (C5/C6) properties (after downstream, OBL naphtha fractionation) Specific gravity

0.67

API

80

RON

75

Paraffin’s

84 vol%

Naphthenes

14 vol%

Benzene

2.0 vol%

Olefins

0.0

Sulfur content

1 ppm 8 PEP REVIEW 98-7

Table 1 (Concluded) PRODUCTS YIELD AND QUALITY

Heavy Naphtha (C7-310°F) Properties Specific Gravity

0.75

API

57

RON

60

Paraffins

48 vol%

Naphthenes

46 vol%

Aromatics

6.0 vol%

Sulfur content

1 ppm

Kerosene 310-550°F (Jet Fuel) properties Specific gravity

0.822

API

40.6

Sulfur content

10 wt ppm

Smoke point

24.5 mm

Aromatics

11.0 wt%

Flash point

125°F

52°C

Diesel 550-650°F properties Specific gravity

0.835

API

38

Sulfur content

10 wt ppm

Aromatics

11.0 wt%

Polyaromatics

0.20 wt%

Cetane Index

67

Flash point

320°F

160°C

9 PEP REVIEW 98-7

PROCESS DISCUSSION AND RATIONALE The design of the HyCycle hydrocracker, the make up hydrogen compression, reaction loop, vapor recovery and liquids recovery, depend on desired product slate. This is an attempt to optimize the design for maximum diesel yield at the expense of kerosene and naphtha. The following issues have been considered while establishing the design basis. Feedstock The feedstock is a blend of atmospheric and vacuum gas oils produced from Arabian Light Crude oil, 680-980°F/BP. The blend has the following distillation curve:

ASTM D1160 Distillation

°F

°C

IBP

611

322

10%

702

372

30%

763

406

50%

825

441

70%

896

480

90%

966

519

EP

1040

560

The 21.5 API gravity oil (0.9246 specific gravity) contained 21 wt% monoaromatics and 32 wt% total aromatics. Sulfur, nitrogen and conardson carbon contents were:

Sulfur

2.5 wt%

Total nitrogen

0.09 wt%

Conardson carbon

0.15 wt%

Hydrocracking catalyst In general Amorphous or zeolitic catalyst is used for middle distillate hydrocracking and both produce high quality product. The amorphous is somewhat more distillate selective, however requires a higher reaction temperature. On the basis of UOP data specific to low-conversion HyCycle, it was decided to use zeolite based catalyst. This rational can be reviewed on a case specific basis. Potential heat recovery at the reactors The estimated heat of reaction is 45-50 Btu/Scf of reacted hydrogen, which amounts to about 105-115 MM Btu/hr heat release in the reactors and end up as low level heat. Since reactors pressure is 1,800-1,900 psig as opposed to 2,400 psig in the more traditional design, the use of an isothermal reactor with 2,000 psig steam generation at 635°F may be worth further 10 PEP REVIEW 98-7

consideration. The steam pressure is above reaction pressure, thus avoiding a concern of hydrocarbon leaks into the steam system, the steam can be let down in pressure if need to, and can be superheated at the fired heaters. About 150,000 lb/hr of steam can be recovered assuming boiler feed water preheat to about 450°F by recycle gas cooling. After the steam is superheated to 930°F its motive power can amount to 20,000 kW. At the same time, much of the lower level heat load on the air coolers will be eliminated. Almost needless to say that under such a scenario, the load on the fired heaters will increase substantially due to the assumed superheating. Conversion Data from UOP’s NPRA public release suggests 20-40% conversion per pass, US 5,980,729 suggests a case with 30% conversion per pass [R98-07-001]. It is understood that the above conversion is in the high pressure hot loop. It is further understood that this low conversion low severity is applied to avoid cracking of diesel product that recycles in the loop. It was judged that 33% conversion per pass is a reasonable optimum in the high pressure loop. It was assumed that the combined conversion per pass including the fractionation, is 83% per long pass. The ultimate conversion(3) at 650°F TBP cut point is assumed at 98.5 wt%. In 2001 issue of NPRA publications [R98-07-001] UOP indicates an ultimate conversion of 99.5%, however the cut point is not indicated in NPRA issue. Additional information received by UOP refers 99.5% conversion at 730°F cut point and the 650°F cut point in this case is the basis of the adjustment in the overall conversion. The conversion is at 650°F cut point, however if kerosene is blended to the diesel pool, higher cut point, gas oil material could be blended into the diesel pool. This may result in higher middle distillate yield. Hydrogen Requirement Hydrogen is consumed by hydro-desulfurization, hydro-denitrification, aromatics saturation and hydrocracking. If VGO from delayed coking or visbreaking, is combined in the feed, an additional amount of hydrogen would be needed for saturating olefins. About 4% of the hydrogen is dissolved in the liquid product and purged from the high pressure loop. The net hydrogen consumption was estimated at 1,630 SCF/BBL feed, about 56% in hydrocracking reactor and 42% in hydrotreating reactor where most of hydrotreating is dedicated to aromatic saturation. The consumption in finishing reactor is estimated at 2% of the hydrogen. The pressure in the flash drum, 300 psig is suitable for hydrogen recovery by PSA (pressure swing adsorption) [see PEP Report 212 “Options for Refinery Hydrogen,” (1994)]. A diesel oriented refinery would normally have a relatively smaller catalytic reforming unit which co-produces hydrogen at about 90% purity. Under the most reasonable scenario a dedicated hydrogen plant, in this case 60 MM Scfd (67,000 M3/hr) at high purity, and less than 100 ppm CO, would be required. Steam reforming of light naphtha, about 6,000 bpsd, or LPG is thought to be worth a serious consideration as a hydrogen source, unless low cost natural gas is available. Under this scenario, the light naphtha and LPG produced in the hydrocracking can satisfy about 80% of the feed to the hydrogen plant. Other light naphtha sources such as straight run naphtha from a crude unit could be a good source for hydrogen production. Any impure hydrogen resulting from catalytic reforming, say about 1,500 Scf/bbl, would be used as a separate hydrogen stream for less severe hydroprocessing services such as straight run diesel hydrotreating at about 1,000 psig. 3x50% make up hydrogen compressors have been assumed, however 2x100% units could be a reasonable selection depending on a very case specific basis. Air cooling combined with trim water cooling was assumed for interstage cooling. However, depending on water quality, air coolers could be avoided, thus resulting in cost reduction. The hydrogen recycle loop is cooled 11 PEP REVIEW 98-7

only by air coolers as a safety concern to avoid any of the high pressure rich gas from leaking to cooling water system. It is reasonable to assume that usage of cooling water could result in lower capital and compression cost, however, at some added cost to overcome safety concerns. All these issues could be addressed on a case specific basis. Reactors Based on prior art and UOP patent US 5,980,729, it was estimated a cold liquid space velocity of 1.4 /hr for hydrocracking and 2.8/hr for hydrotreating where catalyst amounts to 80% of reactor T-T volume. Details of the reactor internals are discussed in [R98-07-001]. A design pressure margin of 10% was allowed. Design margins of 5% could be possible under a different ASME code and these issues could be addressed on a case specific basis. External insulation of the reactors will bring down the outside temperature to 140°F. The option of internal insulation, a “cold wall” design is thought to be of lower capital investment but was judged to present operating and safety risks. The Finishing Reactor/ Enhanced Separator uses [US 5980729] hydrogen heating as a reflux duty. In this particular design, reflux duty can be provided by generating 600 psig steam and hot hydrogen for diesel stripping will come from the reheated hydrogen sources. An alternative concept for stripping could be generating 2,000 psig steam in reactor effluent coolers, or by isothermal reactor, and using the steam for stripping the diesel instead of hydrogen. In this case the reflux duty will be provided by generating 600 psig steam. Acid gas removal The ammonia and some 10% of the H2S is removed by water wash. The lean amine solution, such as MDEA, at reactor loop pressure (1,750 psig) removes the bulk of the H2S. Higher absorption pressures above 300 psig is not advantageous from the standpoint of amine circulation for H2S removal and result in higher capital cost associated with the higher pressure and cost of pumping. As an alternate approach, the H2S could have been purged with the hydrogen rich light gases at 300 psig and to proceed to 300 psig MDEA absorber. The down side of this low pressure approach, would be the build up of H2S in the reactor loop, reducing the partial pressure of hydrogen and increasing hydrogen losses to purge. This issue could be addressed on a case specific basis while examining all other issues such as cracking activity and selectivity that could be somewhat affected by the content of H2S and ammonia in the highpressure reaction loop. Acid gas removal is discussed in Pep Report 216, “Acid Gas Treatment and Sulfur Recovery”, (1997). Recycle compressor A centrifugal compressor, 3,300 kW operating was selected for the recycle loop. Although a centrifugal compressor is of lower cost and higher reliability compared with a reciprocating compressor, several issues need to be sorted out on a case specific basis. These issues are related to variance in conditions between start of run (SOR) and end of run (EOR). Variation in molecular weight, resulted from light gas production, and pressure drop, resulted from changes in catalyst, could be key issues in actual design. A steam turbine drive, using 600 psig/750°F steam, condensing at vacuum, with option of steam extraction at 150 psig will provide a good speed control on the compressor. The steam turbine would handle variance of feed and changes from start of run (SOR) and end of run (EOR). A multi-stage steam turbine will provide over 70% adiabatic efficiency. An electric motor drive could also be a viable selection depending on a case specific and overall steam balance of the refinery. 12 PEP REVIEW 98-7

Product fractionation At this point, we assume a LPG stripper at 120 psig, followed by an atmospheric column, at 10 psig at the receiver. Steam is generated in pump-around loops and kerosene-diesel product side strippers. As an alternate, a combination of atmospheric distillation along with mild vacuum, of 100-150 mm Hg, could be a viable option, depending on a case specific basis. Combined full range naphtha C5-310°F is sent to OBL for further product recovery and upgrading. The assumed downstream naphtha fractionation could become heat integrated with pump-around loops of the main distillate fractionator. Heaters The more common practice for recycle heater is to use a mixed phase heater in the reaction loop for both hydrogen recycle and liquid feed recycle. The alternate common approach would be to use separate heaters for the liquid feed and gas recycles. In the Hycycle case, the hydrotreating and hydrocracking reactors are in reversed arrangement and the UCO recycles hot within the high pressure loop. Because of this configuration, heating of hydrogen recycle as an exclusive source of thermal duty, is the chosen option. This gas recycle heating results in somewhat higher exit temperature and higher skin temperature, compared with mixed phase heating, thus fuel gas firing will be the highly preferred mode of operation. The heater for the fractionation unit is operating in lower severity in both temperature and pressures, thus fuel firing or dual firing would be an acceptable approach. The UCO is purged at 650°F from the recycle seems to be a good low sulfur low nitrogen fuel source. Materials of Construction All the issues of hydrogen graphitization and sulfur corrosion resulting from high pressure, high temperature are fully considered, using the latest version of “Nelson Curves”. For a conservative design the following is suggested. For hydrocracking reactor, we selected 2.25% Cr 1.0 % Mo. For hydrotreating reactor we selected 2.25 Cr 1.0% Mo, with 347 stainless cladding. For the Finishing Reactor and enhanced separation use 1.25% Cr 0.5% Mo and SS 317 cladding. Outside insulation will bring the temperature to 140°F. For reactor loop heater, 2.25 Cr 1.0 Mo is used. For the fractionator heater and all exchangers above 450°F we used 1.25% Cr 0.5% Mo. All services below 450°F use C.S, regardless of pressure. For sulfur corrosion above 625°F use 321 SS in heat exchangers. Equipment size limitation A size limitation of 17' - 0” (5,185 mm) is assumed for transportation of pressure vessels. This limit along with other design considerations set the diameter of the hydrocracking and hydrotreating reactors. A key design consideration is the wall thickness of the reactors. For 35,000 bpsd HyCycle this wall thickness is expected to be in the order of 12” which would represent a near maximum practical fabrication. The weight of the hydrocracking reactor is expected to be in the order of 2,200,000 lb (1,000 ton) which could present a near limit for transportation and erection, especially in emerging markets areas. Products quality The diesel product will meet the Euro IV specification and will be mixed with hydrotreated straight run diesel, about 35,000 BPSD as produced in the crude distillation unit. The kerosene product, as described, meets the specifications for jet fuel. It would be a common practice to mix 13 PEP REVIEW 98-7

this kerosene product with straight run (310-550°F) kerosene material as produced in crude unit after a MEROX (mercaptans oxidation) unit, if needed. The heavy naphtha product is of low octane, thus could be very synergistic for catalytic reforming (CCR) along with straight run naphtha. The light naphtha (C5/C6) could be a good stock for hydrogen production by steam reforming. The LPG (C3/C4) could meet typical market specifications. Waste treatment The hydrocracking process design feature includes water wash of tube sheets and other elements of the cold sections of the reactor loop. This water, about 80 gpm after dissolving ammonium sulfide, is directed to waste water. For the particular case as depicted in this design basis ammonia recovery, being only about 5.0 tpd, would not be economical. The wastewater is directed to sour water stripper along with sour water from other units such as distillate hydrotreating and CCR. The H2S and ammonia are stripped, and routed to conventional Claus type sulfur plants and associated tail gas treating. Ultimately, 99.5% or higher sulfur recovery is achieved. All the ammonia is oxidized, in the sulfur plant, to elemental nitrogen. Additional wastewater, about 65 gpm will result from live steam stripping in fractionation. This water will be routed to API separator with other wastewater streams. PROCESS DESCRIPTION The conceptual design for a 35,000 B/SD gas oil HyCycle hydrocracker is shown in Figure 1. The process design, yields and utility consumption rates are based upon engineering principles, computations, published information and other nonconfidential information. The design may or may not be similar to that licensed or otherwise used. The product yields at the reactor discharge and product quality are listed in Table 1. Table 2 provides the flows and approximate composition of the major streams. Major equipment size and material of construction are listed in Table 3. Estimated utility consumption is summarized in Table 4. The plant consists of two sections: •

Section 100—H2 Compression and Reactor Loop

•

Section 200—Product Recovery

Section 100, H2 compression and reactor loop The hydrogen is produced at 99.9 vol% purity by steam reforming of light naphtha or LPG. A partial source of hydrogen could be reformer (CCR) gas after hydrogen purification. Hydrogen at 340 psig, 100°F (25.0 kg/cm2-a, 38°C) is compressed in K-101 consisting of 3 stages. Average adiabatic efficiency of 83.0% and 97% motor efficiency was assumed. The first stage discharge is 640 psig and achieves a discharge temperature of 225°F (61 kg/cm2-a, 107°C). Second stage discharge is 1,130 psig, 225°F (80.5 kg/cm2-a, 107°C) and 3rd stage discharge is 2,050 psig, 225°F (142.7 kg/cm2-a, 107°C). Hydrogen at 150°F proceeds to mix with the discharge of the recycle gas compressor K-102 at 2,020 psig. In the discharge of stages No-1 and No-2 air cooling (E-101, E-103) is employed down to 150°F (65°C) and trim water cooling (E102, E-104) to 100°F (38°C). Total compression power is 5,060 kw, motor power is 3x2600 Kw and are rated for 3,300 kw each. Cooling water flow rate at 90°F and assumed 20°F temperature rise. VGO/AGO feed from intermediate storage, and 180°F (82°C), VGO from vacuum unit and AGO from crude unit (not shown) are mixed, pumped to 2,200 psig which is the needed delivery 14 PEP REVIEW 98-7

pressure to reactor loop. The feed sent to a filter M-101 of 10 microns cut, prior to being heated from 300°F to 600°F in E-107. For cold feed, exclusively from storage at 180°F, the feed is diverted first to E-106, preheated to 300°F and then proceed to E-107 for final heating to 600°F and prior to mixing with effluent of hydrocracking reactor R-103. In the hydrotreating reactor, all sulfur compounds are converted to H2S, nitrogen compounds are converted to ammonia almost all di and tri aromatics are saturated along with about 55% of the mono-aromatics. Trace organomethalics are removed in the first bed. Mixed phase fluid from R-101 at 730°F is cooled to 640°F in three successive heat exchanger, prior to entering R-102, enhanced separator finishing reactor: Preheating hot UCO recycle at E-108, preheating hydrogen at E-109 and generating 600 psig saturated steam at E-110. The fluid enters R-102 which has three functions: vapor liquid separation in the middle section, stripping the UCO liquid of distillate material, using preheated hydrogen, and residual sulfur removal in the finishing catalytic section. The hydrogen preheating is providing reflux thus minimizing UCO vapor. The finishing reactor bed removes traces of mercaptans that could be formed by recombining of H2S with traces of olefins. Overhead of R102 proceeds to E-107 for feed heating as discussed above. Liquid separated from the bottom of R-102 is reheated in E-108 and recycles to cracking in R-103. In R-103 recycle UCO from product fractionation (area 200) heated hydrogen as well as cold shot of cold hydrogen combined in four beds reaction zone. The desulfurized feed is undergoing 33% cracking to a cut point of 650°F. Hydrocracked mixture at 725°F merges with fresh feed as discussed earlier. Two parallel services are used for hydrogen cooling from E-107. The first is hydrogen preheat at E-112 and the second is low pressure steam generation at E-111. The combined stream from E-113 and E-106 goes through E-114 dedicated as cold feed preheat. If feed is available at temperature below 300°F the flow rate to E-111 is reduced and more heat becomes available to preheat feed at E-106. Portion of UCO from R-102, up to 10%, is let down from loop pressure, 1800 psig to a mid pressure, 320 psig, hot flash drum V-107 this let down control the build up of HPNA and other non-crackable material in the reactor loop. V-107 release most of dissolved hydrogen in the hot UCO this hydrogen is cooled in E-114 and merge with a purge stream drawn from recycle compressor suction drum V-106 this purge stream is free of H2S. The hydrogen rich stream is routed to cold separator V-201 in section 200. Mixed phase of hydrogen and hydrocarbon from E-106 entering air cooler E-113 at about 280°F. A wash water is introduced to wash any ammonium sulfide material from the tubes of E113. Hydrocarbon liquids as well as sour water loaded with all the ammonia and some of the H2S are separated from hydrogen, H2S and light gas in Cold Separator V-105. Overhead from V-105 containing hydrogen, light hydrocarbon gas and H2S proceed to loop amine scrubber C-101 for bulk H2S removal with 50 wt% MDEA. Rich amine solution product is let down to amine regeneration at OBL. Overhead from C-101 at 1750 psig (124 kg/cm2-a) enters the Recycle Gas Compressor K.O Drum V-106 and then to K-102, and is compressed to 2,050 psig. K-102 is a centrifugal compressor driven by a steam turbine. The steam turbine is powered by 600 psig and 750°F (42- kg/cm2-g 400°C) steam, exhausting to 3.5” Hg. Compressor K-102 power is about 3,300 kW. The steam requirement for the turbine is about 35,000 lb/hr, (16 TPH). Recycled gas at about 2,000 psig and 165°F is split about 15 % of it is mixed with compressed make-up hydrogen from K-101 and sent to “cold shot” in the hydrocracking reactor R-103 and also to reflux /stripping gas in the finishing reactor /enhanced separator section of R-102. The balance of the hydrogen is preheated to 625°F in feed effluent exchanger E-112 and E-109 prior to entering to charge heater F-101. The outlet of the charge heater F-101, combined with recycled UCO is the feed to Hydrocracking Reactor R-103 at a controlled temperature of 700°F or any other temperature dictated by catalyst performance and overall heat integration.

15 PEP REVIEW 98-7

Section 200 product recovery Liquid 1,765 psig from V-105 rich in diesel kerosene and naphtha is separated from sour water and let down to Cold Flash Drum V-201 at 300 psig. Dissolved hydrogen containing traces of H2S methane and ethane combines with purge hydrogen and flash hydrogen from area 100. This crude hydrogen stream about 3.0 Mmscfd at about 80 vol % is release to PSA hydrogen recovery at OBL or fuel gas system. Liquids from V-201 depleted of most of the hydrogen and portion of CH4 and C2H6 is let down to C-201, LPG stripper operating at 120 psig (9.4 kg/cm2-a) where essentially all C4 and most C3 are recovered as mixed liquid, about 1,600 BPSD with 38 liquid vol% propane. About 20% of C3 along with residual methane and ethane are released to refinery fuel gas header operating at 80 psig. The stripper feed is preheated by E-201 using 50 psig steam, and reboiled by 600 psig steam at E-203 and stripping is enhanced by small injection of 150 psig live steam. Heavy liquid from Hot Flash Drum V-107 is let down to the coil of fired heater F-201 and partially vaporized at end point of 700 F. Vapor /liquid mixture is introduced to the flash zone of the Fractionation Column C-202 with total of 50 trays. The UCO at 650°F and 30 psig, from the bottom of fractionation column C-202 enters a booster pump provides suction pressure to the recycle pump which recycles the UCO at 2200 psig to the hydrocracking reactor R-103 in area 100. A HPNA purge of about 525 BPSD is drawn from the discharge of the booster pump and could be used as a major source of fuel for F-201. The 10,066 BPSD full range naphtha C5-310 F is recovered as overhead product of C-202 and sent to OBL by the reflux pump. About 14,836 BPSD kerosene product from Kerosene Stripper C-203 is cooled by E-209, generating 50 psig steam , followed by air cooler E-210. The diesel product, about 13,461 BPSD from Diesel Stripper C-204 is cooled by generating 150 psig steam at E-211, followed by air cooler E-212 to 160°F and sent to OBL.

16 PEP REVIEW 98-7

Table 2 Hycycle ™ Unicracking Hydrocracking for Middle Distillate STREAM FLOWS CAPACITY: 1372 MILLION LB/YR

Water Hydrogen Hydrogen Sulfide Ammonia Methane Ethane Propane Butane ( mix) Pentane ( mix) Hexane ( mix) C7-310 F TBP 310-550 F TBP 550 F-650 F TBP 650 F-980 F TBP

Mol. Wt. 18.0 2.0 34.1 14.3 16.0 30.1 44.1 58.1 72.2 86.2

Total, lb/hr Total, kg/hr

Water Hydrogen Hydrogen Sulfide Ammonia Methane Ethane Propane Butane ( mix) Pentane ( mix) Hexane ( mix) C7-310 F TBP 310-550 F TBP 550 F-650 F TBP 650 F-980 F TBP Total, lb/hr Total, kg/hr

Mol. Wt. 18.0 2.0 34.1 14.3 16.0 30.1 44.1 58.1 72.2 86.2

(1)

(2)

(3) 895 77,450 110 -24,265 35,950 19,450 13,385 2,300 600 110 11

13,165 -------------

------------471,625

955,305

13,165 5,972

471,625 213,925

1,129,831 512,483

(9) -Trace ---------Trace Trace 82,000

(10) -Trace ---------Trace Trace Trace

82,000 37,195

Trace Trace

(11)

STREAM FLOWS, LB/HR (4) (5) 895 895 70,885 65,080 110 11,618 -472 24,265 24,415 37,645 37,385 24,980 25,580 21,545 22,345 11,675 12,675 15,510 17,110 73,400 78,400 168,190 178,190 153,182 163,182 999,579 964,514 1,601,861 726,592

1,601,861 726,592

STREAM FLOWS, LB/HR (12) (13) 743 180 65,395 12,705 91 10

(6) 160 12,515 9 -1,985 2,265 1,600 1,095 182 45 10 1 ---

(7) -Trace ---------Trace Trace 947,305

(8) -Trace -----------82,000

19,867 9,012

947,305 429,690

82,000 37,195

(15) 740 54,995 8,176 330 19,820 27,755 19,025 16,410 9,000 12,010 54,955 125,475 114,282 10,500

(16) 41,055 77,995 11,630 472 26,735 39,650 27,180 23,440 12,857 17,155 78,510 179,250 163,260 15,000

473,473 214,764

714,189 323,951

79 106 73 48 5 2 2

22,080 32,558 17,689 12,180 2,099 553 190 68 80 955,305

2,185 2,940 1,760 1,205 200 50 10 1 ---

(14) 1,055 77,995 11,630 472 26,735 39,650 27,180 23,440 12,857 17,155 78,510 179,250 163,285 15,209

541 245

1,109,031 503,048

21,246 9,637

674,423 305,913

10 215 1

----

17 PEP REVIEW 98-7

Table 2 (Continued) Hycycle ™ Unicracking Hydrocracking for Middle Distillate STREAM FLOWS CAPACITY: 1372 MILLION LB/YR

Water Hydrogen Hydrogen Sulfide Ammonia Methane Ethane Propane Butane ( mix) Pentane ( mix) Hexane ( mix) C7-310 F TBP 310-550 F TBP 550 F-650 F TBP 650 F-980 F TBP

Mol. Wt. 18.0 2.0 34.1 14.3 16.0 30.1 44.1 58.1 72.2 86.2

Total, lb/hr Total, kg/hr

Water Hydrogen Hydrogen Sulfide Ammonia Methane Ethane Propane Butane ( mix) Pentane ( mix) Hexane ( mix) C7-310 F TBP 310-550 F TBP 550 F-650 F TBP 650 F-980 F TBP Total, lb/hr Total, kg/hr

Mol. Wt. 18.0 2.0 34.1 14.3 16.0 30.1 44.1 58.1 72.2 86.2

(17) 39,900 5 1,890 472 ------10 10 10 5

(18) Trace 325 550 Trace 1,572 1,781 6,058 8,913 10,371 16,505 78,288 179,170 163,170 14,995

(19) 1,155 77,665 9,190 -24,750 38,343 21,449 14,745 2,511 658 220 70 80 --

42,302 19,188

481,698 218,495

190,836 86,562

(25) 4,000 --------------

(26) 3,900 65 190 -780 1,550 7,900 10,300 1,200 300 -----

(27)

4,000 1,814

26,185 11,877

13,610 6,173

20 20 -5 50 4,785 8,640 80 10 -----

STREAM FLOWS, LB/HR (20) (21) 1,065 1,055 77,665 77,450 110 110 --26,330 26,250 38,343 37,763 21,449 21,049 14,745 14,480 2,511 2,481 658 648 220 210 70 70 80 80 --183,246 83,119

181,646 82,393

STREAM FLOWS, LB/HR (28) (29) 10 3,900 65 160 10 --757 1,431 1,203 1 193 1 5 -2 ---------3,826 1,735

3,911 1,774

(22) 35,000 -------------35,000 15,876

(30)

(23)

(24)

260 350 --

-750 300 50 20

762 1,481 6,008 8,893 10,371 16,505 78,288 179,120 163,170 14,995

1,730 785

479,858 217,660

-------

(31) 100 10

--

65 200

(32) 45

------

30 ------

20 60 10,286 16,493 78,288 179,120 163,170 14,995

30 90 15,430 24,740 117,432 ----

20 60 10,286 16,493 78,288 ----

462,542 209,805

157,767 71,562

105,177 47,707

18 PEP REVIEW 98-7

Table 2 (concluded) Hycycle ™ Unicracking Hydrocracking for Middle Distillate STREAM FLOWS CAPACITY: 1372 MILLION LB/YR

STREAM FLOWS, LB/HR Water Hydrogen Hydrogen Sulfide Ammonia Methane Ethane Propane Butane ( mix) Pentane ( mix) Hexane ( mix) C7-310 F TBP 310-550 F TBP 550 F-650 F TBP 650 F-980 F TBP Total, lb/hr Total, kg/hr

Mol. Wt. 18.0 2.0 34.1 14.3 16.0 30.1 44.1 58.1 72.2 86.2

(33)

(34) 50

(35)

(36)

(37)

50

30

28

----------179,120 ---

163,170 --

------------97,000

------------90,000

2 ------------7,000

179,170 81,270

163,220 74,035

97,030 44,012

90,028 40,836

7,002 3,176

-----------

19 PEP REVIEW 98-7

Table 3 HYDROCRACKING TECHNOLOGY FOR MIDDLE DISTILLATE MAJOR EQUIPMENT CAPACITY: 1,372 MILLION LB/YR (622,000 T/YR) DIESEL FUEL AT 0.96 STREAM FACTOR EQUIPMENT NUMBER -------------------

NAME ----------------------------------------------

SIZE ----------------------------------------

MATERIAL OF CONSTRUCTION REMARKS ---------------------------------------------------------- ----------------------------------------------------------------------------------

REACTORS R-101

HYDROTREATING REACTOR

R-102

FINISHING REACTOR

R-103

HYDROCRACKER

15 50 15 50 15 80

FT DIA FT T-T FT DIA FT T-T FT DIA FT T-T

8 36 10 60 16 120 6 25 7.5 33

FT DIA FT FT DIA FT FT DIA FT FT DIA FT FT DIA FT

CLADDING: 347 SS CLADDING: 347 SS SHELL: 2.25%CR, 1%MO

2.25 CR/1 MO 3 BEDSD 2.25 CR, 1 MO 1 BED; 8 STRIPPING TRAYS 4 BEDS

COLUMNS C-101

AMINE ABSORBER

C-201

LPG STRIPPER

C-202

PRODUCT FRACTIONATOR

C-203

KEROSENE STRIPPER

C-204

DIESEL STRIPPER

SHELL: C.S. TRAYS: C.S. SHELL: C.S. TRAYS: C.S. SHELL: C.S. TRAYS: C.S. SHELL: C.S. TRAYS: C.S. SHELL: C.S. TRAYS: C.S.

15 VALVE TRAYS, 24 INCH SPACING

3 STAGES EACH, RECIPROCAL; MOTOR DRIVE. 2 OPERATING, 1 SPARE. 1 STAGE, CENTRIFUGAL; TURBINE DRIVER

24 VALVE TRAYS, 24 INCH SPACING 50 VALVE TRAYS, 24 INCH SPACING 8 VALVE TRAYS, 24 INCH SPACING 10 VALVE TRAYS, 24 INCH SPACING

COMPRESSORS K-101A-C

H2 MAKE UP COMPRESSORS

3,500 BHP

C.S.

K-102

H2 RECYCLE COMPRESSOR

4,025 BHP

C.S.

HEAT EXCHANGERS E-101A,B E-102A-C E-103A,B E-104A-C E-105A,B E-106 E-107A,B E-108 E-109 E-110 E-111

H2 COMPRESSOR COOLER 1

380 1.7 H2 COMPRESSOR TRIM 1 COOLER 700 1.2 H2 COMPRESSOR COOLER 2 406 1.7 H2 COMPRESSOR TRIM 2 COOLER 600 1.2 H2 COMPRESSOR COOLER 3 406 1.7 H2-FRESH FEED EXCHANGER 5,000 31 FRESH FEED -H2 EXCHANER 7,000 45.5 RECYCLE UCO-EFFLUENT EXCHANGER 10,000 50 H2 FEED-EFFLUENT EXCHANGER 5,000 45 R-101 EFFLUENT-STEAM EXCHANGER 2,000 37 R-102 CONDENSER 1,500 21

SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR

SHELL: C.S. TUBES: C.S. SHELL: C.S. TUBES: C.S. SHELL: C.S. TUBES: C.S. SHELL: C.S. TUBES: C.S. SHELL: C.S. TUBES: C.S. SHELL: C.S. TUBES: C.S. SHELL: 1.25CR, 0.5 MO TUBES: 1.25 CR, 0.5 MO SHELL: 2.25CR, 1 MO TUBES: 321 SS SHELL: 1.25CR, 0.5 MO TUBES: 321 SS SHELL: C.S. TUBES: 321 SS SHELL: C.S. TUBES: 321 SS

AIR COOLER 1 SPARE IN PARALLEL AIR COOLER 1 SPARE IN PARALLEL AIR COOLER

20 PEP REVIEW 98-7

Table 3 (Continued) HYDROCRACKING TECHNOLOGY FOR MIDDLE DISTILLATE MAJOR EQUIPMENT CAPACITY: 1,372 MILLION LB/YR (622,000 T/YR) DIESEL FUEL AT 0.96 STREAM FACTOR EQUIPMENT NUMBER -------------------

NAME ----------------------------------------------

SIZE ----------------------------------------

E-112

H2-STEAM BOILER

E-113A-D

H2-H2 EXCHANGER

E-114A-C

RECYCLE GAS COOLER

E-115

FLASH H2 COOLER

E-116

TURBINE EXHAUST CONDENSER

E-201

COLD SEPARATOR HTR

E-202

STRIPPER CONDENSER

E-203

STRIPPER REBOILER

E-204

UCO PROD COOLER

E-205

KEROSENE PUMP AROUND COOLER

E-206

DIESEL PUMP AROUND COOLER

E-207A-D

FRACTIONATOR CONDENSER

E-208

NAPHTHA PROD COOLER

E-209

KEROSENE PROD COOLER

E-210A,B

KEROSENE COOLER

E-211

DIESEL PROD COOLER

E-212A,B

DIESEL COOLER

4,500 39 6,000 16.4 6,000 33.3 250 1.5 6,500 35 7,000 30 8,000 27 3,500 27 1,000 13.7 2,200 20 2,000 20 2,800 15 1,500 2.2 1,700 10.7 1,200 8 2,000 13.7 2,229 18

SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR SQ FT MMBTU/HR

MATERIAL OF CONSTRUCTION ---------------------------------------------------------SHELL: C.S. TUBES: C.S. SHELL: C.S. TUBES: C.S. SHELL: C.S. TUBES: C.S. SHELL: C.S. TUBES: C.S. SHELL: C.S. TUBES: C.S. SHELL: C.S. TUBES: C.S. SHELL: C.S. TUBES: C.S. SHELL: C.S. TUBES: C.S. SHELL: C.S. TUBES: C.S. SHELL: C.S. TUBES: C.S. SHELL: C.S. TUBES: C.S. SHELL: C.S. TUBES: C.S. SHELL: C.S. TUBES: C.S. SHELL: C.S. TUBES: C.S. SHELL: C.S. TUBES: C.S. SHELL: C.S. TUBES: C.S. SHELL: C.S. TUBES: C.S.

REMARKS -------------------------------------------------------------------------------

AIR COOLER

AIR COOLER

AIR COOLER

AIR COOLER

PROCESS FURNACES F-101 F-201

FEED-RECYCLE FURNACE PRODUCT FRACTIONATOR

33 MMBTU/HR 161 MMBTU/HR

347 SS CHROME-MOLY

GAS FIRED

C.S. C.S. C.S. C.S. C.S.

3 DAYS SUPPLY 3 DAYS PRODUCTION 3 DAYS PRODUCTION 3 DAYS PRODUCTION 5 DAYS PRODUCTION

TANKS (OFFSITE) T-101A,B T-201 T-202 T-203 T-204

STARTUP OIL TANK NAPHTHA KEROSENE-JET FUEL DIESEL FUEL FUEL OIL

2,205,000 1,250,000 2,000,000 2,000,000 110,000

GAL GAL GAL GAL GAL

21 PEP REVIEW 98-7

Table 3 (Concluded) HYDROCRACKING TECHNOLOGY FOR MIDDLE DISTILLATE MAJOR EQUIPMENT CAPACITY: 1,372 MILLION LB/YR (622,000 T/YR) DIESEL FUEL AT 0.96 STREAM FACTOR EQUIPMENT NUMBER -------------------

NAME ----------------------------------------------

SIZE ----------------------------------------

MATERIAL OF CONSTRUCTION --------------------------------------------------------

REMARKS -------------------------------------------------------------------------------

PRESSURE VESSELS V-101 V-102 V-103 V-104 V-105 V-106 V-107

H2 COMPRESSOR K.O. DRUM 1 H2 COMPRESSOR K.O. DRUM 2 H2 COMPRESSOR K.O. DRUM 3 FEED SURGE DRUM COLD SEPARATOR RECYCLE COMPRESSOR KO. DRUM HOT FLASH DRUM

700 700 700 9,000 32,000 3,400 4,000

GAL GAL GAL GAL GAL GAL GAL

V-201 V-202 V-203

COLD FLASH DRUM LPG REFLUX DRUM FRACTIONATOR REFLUX DRUM

28,500 GAL 1,000 GAL 5,400 GAL

C.S. C.S. C.S. C.S. C.S. C.S. 1.25%CR, 0.5%MO

HORIZONTAL

C.S. C.S. C.S.

HORIZONTAL

CARBON STEEL CARBON STEEL

100 MESH SS SINTERED, BACKFLUSH 100 MESH SS SINTERED, BACKFLUSH

MISCELLANEOUS EQUIPMENT M-101 M-201

FRESH FEED FILTER SYSTEM UCO FILTER SYSTEM

SPECIAL EQUIPMENT S-101

CATALYST

INITIAL CHARGE

PUMPS SECTION --------------100 200

OPERATING ------------------19 10

SPARES -------------9 8

OPERATING BHP ------------------------3,578 1,010

22 PEP REVIEW 98-7

Table 4 HYDROCRACKING TECHNOLOGY FOR MIDDLE DISTILLATE UTILITIES SUMMARY CAPACITY: 1,372 MILLION LB/YR (622,000 T/YR) DIESEL FUEL AT 0.96 STREAM FACTOR UNITS --------------AVERAGE CONSUMPTIONS COOLING WATER PROCESS WATER ELECTRICITY NATURAL GAS STEAM, 50 PSIG STEAM, 150 PSIG STEAM, 600 PSIG STEAM, 50 PSIG STEAM, 150 PSIG STEAM, 600 PSIG

GPM GPM KW MM BTU/HR M LB/HR M LB/HR M LB/HR M LB/HR M LB/HR M LB/HR

PEAK DEMANDS COOLING WATER PROCESS WATER STEAM, 50 PSIG STEAM, 150 PSIG STEAM, 600 PSIG

GPM GPM M LB/HR M LB/HR M LB/HR

BATTERY LIMITS TOTAL ----------------------

SECTION 100 ---------------

SECTION 200 ---------------

6,660 80 10,574 205 28 4 72 -102 -43 -44

3,740 80 9,394 44 --35 -40 --38

2,920 -1,181 161 28 4 37 -62 -43 -6

7,992 96 34 5 86

4,488 96 --42

3,504 -34 5 44

23 PEP REVIEW 98-7

COST ESTIMATES The capital investment and production costs for hydrocracking as oil produced from Arabian Light crude oil for the production of jet and diesel fuels are presented. Costs are based on an U.S. Gulf Coast location for the second quarter of 2000. Overnight construction is assumed (ie, the cost of funds spent during construction are excluded). Investment Costs Table 5 sets forth the total fixed capital (TFC) investment for a 35,000 B/SD HyCycle hydrocracking unit. The costs are calculated at a PEP Cost Index of 624. The design has not been fully optimized either for equipment size or energy consumption. Owner’s costs such as cleared land are excluded. Licensing or royalty fees are regarded as proprietary and are not included in these costs. The boundary limit installed costs was $151 million before contingency. This cost is 14% lower than the corresponding cost of the conventional VGO middle distillate hydrocracker evaluated in PEP Report 211. Both units cracked Arabian Light gas oil. The conventional unit operates at 2645 psi compared to 1910 psig in the HyCycle process. The lower pressure reduces the reactor cost and hydrogen compressor cost. UOP has estimated a 5% capital reduction compared to “previous offerings” due to these factors [R98-007-001]. Table 6 shows the investment costs broken down by section. The major items are the reactors and initial catalyst charge and the hydrogen compressors. The estimated cost of catalysts, about $10 MM, is based on generic industry data from hydrocracking projects. Because of the large size hydrocracking reactors and the subsequent portion of capital cost attributed to these reactors, it should be noted that the process specifications of these items could be highly influenced by site specific factors such as transportation local construction capability. Material of construction design codes, desired run length on the catalyst could be subject to operator’s preference and would affect capital investment. Adding 15 or 25% contingency to the current estimate results in a boundary limit investment of 174 million $ or 189 million $ respectively. Offsite investment before contingency (Table 5) totals 21 million $. The offsite investment is relatively small percentage of the boundary limit investment due to the high cost of the reactors and credit for steam generation within the boundary limits. The cost of the offsite is based on an assumed generic integration with other refinery units such as crude distillation, vacuum distillation sulfur recovery, waste water treating, spent catalyst, and intermediate storage. Offsite investment includes the incremental cost of utilities and storage. Gas oil feed and product storage are based on 3 days production (except fuel oil, 5 days). The allowance for general service facilities includes a control room, roads, fire protection and equipment. The cost of rich MDEA stripping is included in the waste treatment investment. The allocation also covers the flare and relief system and incremental additions to the water collection and treatment system. Contingencies of 15 or 25% bring the total offsite investment to 23.8 million $ or 26.5 million $ respectively. The total fixed capital cost or boundary limits investment is 198 million $ including a 15% contingency or 216 million $ when including a 25% contingency. Production Costs Table 7 lists the unit cost and consumption factors used in estimating the value of feedstocks, products and utilities included in the production cost. By-products are shown as a credit (negative). Feedstock and product values are mid-2000 values determined by PEP. The 24 PEP REVIEW 98-7

gas oil feed is based on 10 vol% atmospheric gas oil and 90% vacuum gas oil. The on stream factor is 0.96. Raw material cost of 24.51 ¢/lb of diesel fuel includes catalyst replacement every three years (UOP expects 4 years catalyst life with their new catalysts [R97-07-001]. By-product sales and credits of mainly jet fuel and full range naphtha offset most of the raw material cost. Utility costs, 0.52 ¢/lb of diesel, are relatively low due to credit for steam generation. Plant cash costs total 4.48 ¢/lb of diesel fuel. Adding depreciation at 10%/yr of total fixed capital of $216 million (including a 25% contingency) and an allowance for general administration, sales and research expenses gives a net production cost of 6.15 ¢/lb. At a 25% contingency factor, the product value is 9.51 ¢/lb of diesel fuel. Product value of units of 0.5 and 1.5 times the base case capacity were also calculated at a 0.96 on stream factor and a constant 25% ROI. With a 25% contingency, the product value from the half capacity plant is 11.49 ¢/lb, 14% higher than the 10.08 ¢/lb of the base case. Increasing capacity to 150% of the base case reduces the product value 4.6% to 9.62 ¢/lb. With a 15% contingency, the respective product values at 0.5 and 1.5 times the base case capacity are 10.81 and 9.07 ¢/lb of diesel fuel. With 25% contingency, reducing the time on stream factor to 0.92 from 0.96 reduces the production of diesel fuel to 1315 million lb/yr (12,342 B/D) from 1,372 million lb/yr (12,878B/D). Plant cash costs increase to 4.55 ¢/lb from 4.48 ¢/lb of diesel. Net production cost becomes 6.29 ¢/lb, up from 6.15 ¢/lb. The total product value at 25% ROI rises to 10.39 ¢/lb from 10.08 ¢/lb. Similarly, with 15% contingency, the base unit’s total production cost decreases to 10.21 ¢/lb of diesel with less time on steam. Profitability Based upon a total fixed capital cost (included 25% contingency) of $216 million and a 0.96 operating factor, the total return on investment (ROI) before income taxes, was an attractive 29.3% at the mid 2000 prevailing diesel fuel market value of 10.77 ¢/lb. When the plant capacity is half the base case capacity, the ROI reduces to 21.1%. Increasing capacity to 1.5 times the base case increases the ROI to 32.7%. Similarly, but with a 15% capital cost contingency, the base case unit’s ROI increases to a very attractive 33.6%, the half capacity case returns 24.7% and the 1.5 times case return rises to 37.4%. At a 0.92 operating factor instead of 0.96 and a product value of 10.77¢/lb of diesel fuel, the base case unit’s ROI decreases 2.1 percentage points to 27.2% with a 25% contingency and rises 2.0 points to 31.4% with a 15% contingency. At 25% contingency, the 0.5 and 1.5 times capacity plants' returns are 19.2% and 30.5% respectively. Also with a 15% contingency, the 0.5 and 1.85 times plant capacity plants returns are 22.7% and 34.9% respectively, down 2.0 and 2.5 percentage points when the time on stream is reduced to 0.92.

25 PEP REVIEW 98-7

Table 5 HYDROCRACKING TECHNOLOGY FOR MIDDLE DISTILLATE TOTAL CAPITAL INVESTMENT CAPACITY: 1,372 MILLION LB/YR (622,000 T/YR) DIESEL FUEL AT 0.96 STREAM FACTOR PEP COST INDEX: 624

COST ($1,000) -----------BATTERY LIMITS EQUIPMENT, F.O.B. REACTORS COLUMNS VESSELS & TANKS HEAT EXCHANGERS FURNACES COMPRESSORS SPECIAL EQUIPMENT MISCELLANEOUS EQUIPMENT PUMPS TOTAL DIRECT INSTALLATION COSTS INDIRECT COSTS UNSCHEDULED EQUIPMENT, 10% BATTERY LIMITS, INSTALLED CONTINGENCY, 25% BATTERY LIMITS INVESTMENT OFF-SITES, INSTALLED CLARIFIED WATER COOLING WATER PROCESS WATER BOILER FEED WATER STEAM TANKAGE UTILITIES & STORAGE GENERAL SERVICE FACILITIES WASTE TREATMENT TOTAL CONTINGENCY, 25% OFF-SITES INVESTMENT TOTAL FIXED CAPITAL

CAPACITY EXPONENT -------------------------UP DOWN ---------- ----------

24,383 1,373 1,834 5,666 5,897 9,921 10,000 700 2,670 --------62,444

0.98 0.95 0.84 0.78 0.82 0.76 1.00 1.00 0.79

0.90 0.81 0.97 0.63 0.82 0.76 1.00 1.00 0.74

0.90

0.85

37,163 37,915 13,752 --------151,274

0.86 0.84 0.87

0.73 0.76 0.79

0.87

0.79

37,819 --------189,093

0.87

0.79

0.87

0.79

578 1,284 302 2,695 2,739 6,662 --------14,261

0.74 0.92 0.62 0.52 0.48 0.65

0.64 0.92 0.62 0.39 0.00 0.65

0.62

0.47

4,067 2,888 --------21,216

0.86 0.84

0.77 0.74

0.70

0.56

5,304 --------26,519

0.70

0.56

0.70

0.56

215,612

0.85

0.76

26 PEP REVIEW 98-7

Table 6 HYDROCRACKING TECHNOLOGY FOR MIDDLE DISTILLATE CAPITAL INVESTMENT BY SECTION CAPACITY: 1,372 MILLION LB/YR (622,000 T/YR) DIESEL FUEL AT 0.96 STREAM FACTOR PEP COST INDEX: 624 REACTOR LOOP PRODUCT RECOVERY ------------------------------------------------- ------------------------------------------------CAPACITY CAPACITY EXPONENT EXPONENT COST ------------------------COST ------------------------($1,000) UP DOWN ($1,000) UP DOWN ---------------------- ------------------------------- ---------BATTERY LIMITS EQUIPMENT, F.O.B. REACTORS COLUMNS VESSELS & TANKS HEAT EXCHANGERS FURNACES COMPRESSORS SPECIAL EQUIPMENT MISCELLANEOUS EQUIPMENT PUMPS TOTAL DIRECT INSTALLATION COSTS INDIRECT COSTS UNSCHEDULED EQUIPMENT, 10% BATTERY LIMITS INSTALLED CONTINGENCY, 25% BATTERY LIMITS INVESTMENT OFFSITES, INSTALLED CLARIFIED WATER COOLING WATER PROCESS WATER BOILER FEED WATER STEAM TANKAGE UTILITIES & STORAGE

24,383 520 1,565 4,466 2,813 9,921 10,000 450 1,707 --------55,824

0.98 1.12 0.86 0.79 0.82 0.76 1.00 1.00 0.83

0.90 0.95 1.03 0.64 0.82 0.76 1.00 1.00 0.76

0.92

0.87

30,654 32,917 11,940 --------131,335

0.87 0.86 0.89

0.77 0.79 0.82

0.89

0.82

32,834 --------164,169

0.89

0.82

0.89

0.82

0.74 0.92 0.62 0.51 0.48 0.65

0.64 0.92 0.62 0.34 0.00 0.65

0.66

0.58

309 721 130 1,028 258 2,772 --------5,219

-853 269 1,200 3,085 --250 963 --------6,620

-0.84 0.75 0.77 0.82 --1.00 0.73

-0.74 0.68 0.62 0.82 --1.00 0.70

0.80

0.75

6,509 4,997 1,813 --------19,939

0.77 0.75 0.78

0.56 0.61 0.64

0.78

0.64

4,985 --------24,924

0.78

0.64

0.78

0.64

0.74 0.92 0.62 0.53 0.48 0.65

0.64 0.92 0.62 0.42 0.00 0.65

0.60

0.41

269 563 102 1,667 2,481 3,890 --------8,972

27 PEP REVIEW 98-7

Table 7 HYDROCRACKING TECHNOLOGY FOR MIDDLE DISTILLATE PRODUCTION COSTS PEP COST INDEX: 624 VARIABLE COSTS UNIT COST ------------------------RAW MATERIALS FRESH FEED 10% AGO HYDROGEN 99.9% CATALYST & CHEMICALS MDEA

CONSUMPTION PER LB --------------------------

¢/LB ----------

6.442 ¢/LB 73.21 ¢/LB 834 ¢/LB 183 ¢/LB

2.89017 LB 0.07717 LB 0.000292 LB 0.000002 LB

18.62 5.65 0.24 NEGL --------24.51

7.806 ¢/LB 11.11 ¢/LB 12.05 ¢/LB 11.55 ¢/LB 13.13 ¢/LB 11.12 ¢/LB 7.671 ¢/LB

-0.01012 LB -0.01156 LB -0.03757 LB -0.05491 LB -0.64451 LB -1.09191 LB -0.04335 LB

-0.08 -0.13 -0.45 -0.63 -8.46 -12.14 -0.33 ---------22.22

GROSS RAW MATERIALS BY-PRODUCTS C1 FUEL GAS C2 FUEL GAS C3 FUEL GAS C4 FUEL GAS FULL NAPHTHA KEROSENE(JET FUEL) H. FUEL OIL TOTAL BY-PRODUCTS

UNIT COST ------------------------UTILITIES COOLING WATER PROCESS WATER STEAM, 600 PSIG STEAM, 50 PSIG STEAM, 150 PSIG ELECTRICITY NATURAL GAS

CONSUMPTION CONSUMPTION PER LB PER KG -------------------------- --------------------------

7.45 ¢/MGAL 1.101 $/MGAL 5.72 $/MLB 3.5 $/MLB 4.51 $/MLB 4 ¢/KWH 3.24 $/MMBTU

2.45 GAL 0.0294 GAL 0.172 LB -0.453 LB -0.239 LB 0.0648 KWH 1,260 BTU

TOTAL UTILITIES

20.4 LITERS 0.245 LITERS 0.172 KG -0.453 KG -0.239 KG 0.143 KWH 698 KCAL

0.02 NEGL 0.10 -0.16 -0.11 0.26 0.41 --------0.52

28 PEP REVIEW 98-7

Table 7 (concluded) HYDROCRACKING TECHNOLOGY FOR MIDDLE DISTILLATE PRODUCTION COSTS PEP COST INDEX: 624 CAPACITY (MILLION LB/YR)* INVESTMENT ($ MILLIONS) BATTERY LIMITS (BLI) OFFSITES TOTAL FIXED CAPITAL (TFC)

686 ------------

1,372# ------------

2,058 ------------

109.0 18.0 --------127.0

189.1 26.5 --------215.6

269.6 35.2 --------304.8

SCALING EXPONENTS

0.76

PRODUCTION COSTS (¢/LB) RAW MATERIALS BY-PRODUCTS UTILITIES VARIABLE COSTS OPERATING LABOR, 4/SHIFT, $33.58/HR MAINTENANCE LABOR, 3%/YR OF BLI CONTROL LAB LABOR, 20% OF OPER LABOR LABOR COSTS MAINTENANCE MATERIALS, 3%/YR OF BLI OPERATING SUPPLIES, 10% OF OPER LABOR TOTAL DIRECT COSTS PLANT OVERHEAD, 80% OF LABOR COSTS TAXES AND INSURANCE, 2%/YR OF TFC PLANT CASH COSTS DEPRECIATION, 10%/YR OF TFC PLANT GATE COSTS G&A, SALES, RESEARCH NET PRODUCTION COST ROI BEFORE TAXES, 29.3%/YR OF TFC PRODUCT VALUE

0.85

24.51 -22.22 0.52 --------2.81

24.51 -22.22 0.52 --------2.81

24.51 -22.22 0.52 --------2.81

0.17 0.48 0.03 --------0.68

0.09 0.41 0.02 --------0.52

0.06 0.39 0.01 --------0.46

0.48 0.02 --------3.99

0.41 0.01 --------3.75

0.39 0.01 --------3.67

0.54 0.37 --------4.90

0.42 0.31 --------4.48

0.37 0.30 --------4.34

1.85 --------6.75

1.57 --------6.05

1.48 --------5.82

0.12 --------6.87

0.11 --------6.16

0.10 --------5.92

5.42 --------12.29

4.60 --------10.76

4.34 --------10.26

----------------------------------* OF DIESEL FUEL # BASE CASE

29 PEP REVIEW 98-7

CONCLUSION If the high yield of diesel product would represent a desired product slate, the preliminary conclusion about the HyCycle(TM) Unicracking technology is very positive. This preliminary conclusion could be enforced by confirmation of the assumption made during the course of this evaluation, in terms of reactor operating conditions, vapor product separation conditions, and product yield. It is recognized that a non-disclosure agreement with UOP would have been required in order to further establish the technical integrity of the above report. FOOTNOTES (1) Cetane Index. A Nomograph for calculating Cetane Index is presented in Standard Methods For Analysis and Testing of Petroleum and Related Products. 1988 issue British Institute of Petroleum. Vol 1 218-3. API gravity and 50% boiling point by ASTM distillation is needed. (2) Per pass conversion: 100- 650°F+ material in distillate product 650°F+ in (fresh feed + recycle) (3) Overall conversion : 100- 650°F+ in distillate product 650°F+ material in feed

(liquid volume)

(liquid volume)

30 PEP REVIEW 98-7

CITED REFERENCES Literature R98-07-001

Antos, G.J., et al., “Unicracking TM Innovations Delivery Profit,” National Petrochemical and Refiners Association 2001 Annual Meeting, New Orleans, LA, (March 18-20, 2001), Paper AM-01-30

R98-07-002

Lamourelle, A.P., et al., “Clean Fuels: Route to Low Sulfur Low Aromatic Diesel,” National Petrochemical and Refiners Association 2001 Annual Meeting, New Orleans, LA, (March 18-20, 2001), Paper AM-01-28

R98-07-003

Adler, K., “2001 In Review: Europe, Asia See Trends of Fuels Regulations,” World Refining, 11,1 (January/February 2001), 36, 38-40

R98-07-004

Peckham, J., “Euro Commission Proposes 10 ppm ULSD Phase-in from 2007,” Diesel Fuel News, 5, 5 (March 5, 2001), 11\

R98-07-005

Roj, A., “Fuel Quality for the Future—an Automotive Industry View,” Cleaner Fuels for Europe workshop, Helsinki, Finland, (Nov. 23-23, 2000)

R98-07-006

U.S. Environmental Protection Agency, “EPA Gives the Green Light on DieselSulfur Rule,” News Release (R-30), (February 28, 2001)

R98-07-007

National Petrochemical and Refiners Association, “NPRA to Challenge Diesel Sulfur Rule in Court: Decision Motivated by Supply Concerns,” News Release (January 23, 2001)

R98-07-008

National Petrochemical and Refiners Association, “NPRA Files Petition for Review of EPA’s Diesel Sulfur Rule,” News Release (February 2, 2001)

R98-07-009

National Petrochemical and Refiners Association, “NPRA Reaction to Reports of Administration Decision on Diesel Rule,” News Release (February 28, 2001)

Patents US 5885440

Hoehn, R.K., et al., (to UOP LLC), “Hydrocracking process with integrated effluent hydrotreating zone,” U.S. 5,885,440 (March 23, 1999)

US 5980729

Kalnes, T.N., (to UOP LLC),”Hydrocracking Process,” U.S. 5,980,729 (Nov. 9, 1999)

WO 97/38066

Cash, I. R., (to Chevron U.S.A.), “Process for Reverse Staging in Hydroprocessing Reactor Systems,” WO 97/38066 (application) (Oct. 16, 1997)

Reports 211

Chang, E.J., “Hydrocracking,” PEP Report 211, SRI Consulting, Menlo Park, CA (April 1994)

212

Leiby, S.M., “Options for Refinery Hydrogen,” PEP Report 212, SRI Consulting, Menlo Park, CA (February 1994)

216

Ma, J.J. “Acid Gas Treatment and Sulfur Recovery,” PEP Report 216, SRI Consulting, Menlo Park, CA (November 1997) 31 PEP REVIEW 98-7

228

Nielsen, R. H., “Refinery Residue Upgrading,” PEP Report 228, SRI Consulting, Menlo Park, CA (May 2000)

32 PEP REVIEW 98-7

Figure 1 (Sheet 1 of 2) (TM)

HYCYCLE UNICRACKING HYDROCRACKING FOR MIDDLE DISTILLATE H2 COMPRESSION AND REACTOR LOOP O

225 F 640 psig

150OF

225OF 1130 psig

O

100 F

150OF

V-102

V-101 E-101

E-102

150OF

V-103

E-104

E-103

600 psig Steam

225OF 2050 psig

100OF

22

E-105

Make-up H2

K-102 1

K-101B

K-101A

K-101C E-116 Vac.

Vent

Wash Water

2020 psig 165OF

V-104

CW

140OF 16

V-105

Atmospheric Gas Oil

21

E-114

Vacuum Gas Oil From Storage

To V-201

18

300OF

V-106 M-101

19

17

2

330OF

E-106

320OF

Waste Water

E-107

600OF

130OF 1750 psig

E-113 20 O

490OF

E-112

4

15

700OF 1910 psig

BFW 690OF

14

Lean MDEA

O

630 F

R-101

13

R-102

R-103

E-111

E-110

3

130 F C-101

Rich MDEA 150OF

O

500 F

640OF

11

6

To V-201

E-115 625OF

O

730 F 1820 psig

Unconverted Oil from Section 200

1800 psig O

O

5

BFW 250OF

E-109

400OF

695 F E-108

BFW

10

V-107

7 8

320OF 1860 psig

12

715OF F-101 1950 psig 36

V-101 H2Co mpressor K.O.1

Unconverted Oil to E-204

9

R-101 F-101 Hydrotreating Feed-Recycle Reactor Furnace

V-104 Feed Surge

M-101 A,B Fresh Feed Filter System K-101A,B,C H2 Make up

PEP Review 98-7

725 F 1860 psig

O

625 F

R-102 Finishing Reactor

V-102 R-103 V-103 H2Co mpressor Hydrocracker H2Co mpressor K.O.2 K.O.3

V-105 Cold Separator

V-107 Hot Flash Drum

V-106 Recycle Compressor K.O.

C-101 Amine Absorber

K-102 H2 Recycle

2001 33

Figure 1 (Sheet 2 of 2) (TM)

HYCYCLE UNICRACKING HYDROCRACKING FOR MIDDLE DISTILLATE PRODUCT RECOVERY

Fuel Gas

E-207 265OF Fuel Gas

H2

180OF

From V-106 & V-107

Steam 10 psig 20,000 lb/hr

28

10 psig 140OF

C-202 BFW

V-202

32

E-205

From V-105

27 18

LPG

Steam 50 psig 10,000 lb/hr

BFW

Water

29

Kerosene 475OF

E-206

Steam 50 psig 32,000 lb/hr

Steam 150 psig

280 F

550 F

Steam 600 psig /750OF 9000 lb/hr 400OF BFW from F-101

25

120 psig

360OF

C-204

O

24

37,000 lb/hr 600 psig Steam

E-203

33

E-209

550OF O

Naphtha

450OF

Steam 150 psig 20,000 lb/hr

11

CW E-208

C-203

450OF

120 psig

E-201

28,000 lb/hr Waste Water

375OF

C-201

V-201

140 F 300 psig

100OF

CW

26

23 O

E-202

V-203

E-210

140OF

BFW

Steam 50 psig 14,000 lb/hr

Steam 150 psig 6000 lb/hr

Diesel

F-201

34

E-211 BFW

E-212

160OF

18,000 lb/hr Steam 50 psig 28,000 lb/hr 30

650OF, 30 psig O

700 F E-204

From V-107 625OF

35

9

To R-103

Fuel Oil

V-201 Cold Flash Drum

PEP Review 98-7

UCO to Fuel Oil

650OF 2200 psig

M-201 36

C-201 LPG Stripper

V-202 LPG Reflux Drum

F-201 Fractionator Feed Heater

M-201 UCO Filter System

C-202 Product Fractionator

37

V-203 Fractionator Reflux

C-203 Kerosene Stripper

C-204 Diesel Stripper

2001 34