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Refinery Process

Refinery process 2019

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Hydrocarbon Processing’s Reﬁning Processes 2006 Handbook reﬂects

the dynamic advancements now available in licensed process technologies, cata-

lysts and equipment. The refining industry is under tremendous pressure to pro-
cess “cleaner” transportation fuels with varying specifications for a global market.
Refiners must balance capital investment and operating strategies that provide the
optimum profitability for their organization. Hence, refining organizations will apply
leading-edge technology in conjunction with “best practices” for refining fuels and
petrochemical feedstocks from crude oil.
HP’s Process Handbooks are inclusive catalogs of established and emerging
refining technologies that can be applied to existing and grassroots facilities.
Economic stresses drive efforts to conserve energy consumption, minimize waste,
improve product qualities and, most importantly, increase yields and throughput.
In further expansion, the process entries presented an expanded description of
the licensed technology including a process flow diagram, product description,
economic information and other vital information. Specific processing opera-
tions to be emphasized include alkylation, coking, (crude) distillation, catalytic
cracking (fluid and resid), hydrocracking, hydrotreating, hydrogen, isomerization,
desulfurization, lube treating, visbreaking, etc.
To maintain as complete a listing as possible, the Refining Processes 2006
Handbook is available on CD-ROM and at our website for paid subscribers.
Additional copies may be ordered from our website.
Photo: The Valero Port Arthur, Texas, Refinery.

Sponsored by:
Technology Solutions
Processes index - 1 [next page]
Alkylation Desulfurization Hydrocracking (ISOCRACKING)
Alkylation, catalytic Dewaxing Hydrocracking (LC-FINING)
Alkylation--feed preparation Dewaxing/wax deoiling Hydrocracking-residue
Alkylation-HF Diesel-ultra-low-sulfur diesel (ULSD) Hydrodearmatization
Alkylation-sulfuric acid Diesel-upgrading Hydrofinishing
Aromatics Ethers Hydrofinishing/hydrotreating
Aromatics extractive distillation Ethers-ETBE Hydrogen
Aromatics recovery Ethers-MTBE Hydrogenation
Benzene reduction Flue gas denitrification Hydrogen-HTCT and HTCR twin plants
Biodiesel Flue gas desulfurization-SNOX Hydrogen-HTER-p
Catalytic dewaxing Fluid catalytic cracking Hydrogen-methanol-to-shift
Catalytic reforming Fluid catalytic cracking-pretreatment Hydrogen-recovery
Coking Gas treating-H2S removal Hydrogen-steam reforming
Coking, fluid Gasification Hydrogen-steam-methane
Coking,flexi Gasoline desulfurization reforming (SMR)
Crude distillation Gasoline desulfurization, ultra deep Hydroprocessing, residue
Crude topping units H2S and SWS gas conversion Hydroprocessing, ULSD
Deasphalting H2S removal Hydrotreating
Deep catalytic cracking Hydroconversion-VGO & DAO Hydrotreating (ISOTREATING)
Deep thermal conversion Hydrocracking Hydrotreating diesel

Sponsored by:
Technology Solutions
Processes index - 2 [previous page]
Hydrotreating/desulfurization Olefins-butenes extractive distillation Treating-gases
Hydrotreating-aromatic saturation Olefins-dehydrogenation of Treating-gasoline and LPG
Hydrotreating-lube and wax light parraffins to olefins Treating-gasoline desulfurization,
Hydrotreating-RDS/VRDS/UFR/OCR Oligomerization-C3/C4 cuts ultra deep
Hydrotreating-resid Oligomerization-polynaphtha Treating-gasoline sweetening
Hydrotreating-residue Paraxylene Treating-kerosine and heavy naphtha
Prereforming with feed ultra purification sweetening
Isomerization
Pressure swing adsorption-rapid cycle Treating-phenolic caustic
Isooctene/isooctane
Refinery offgas-purification and Treating-pressure swing adsorption
Lube and wax processing
olefins recovery Treating-propane
Lube extraction
Resid catalytic cracking Treating-reformer products
Lube hydrotreating
Slack wax deoiling Treating-spent caustic deep neutralization
Lube oil refining, spent
SO2 removal, regenerative Vacuum distillation
Lube treating
Sour gas treatment Visbreaking
Mercaptan removal
Spent acid regneration Wax hydrotreating
NOx abatement
Spent lube oil re-refining Wet gas scrubbing
NOx reduction, low-temperature
Sulfur processing Wet Scrubbing system, EDV
O2 enrichment for Claus units
Sulfur recovery White oil and wax hydrotreating
O2 enrichment for FCC units
Thermal gasoil
Olefin etherfication
Treating jet fuel/kerosine
Olefins recovery

Sponsored by:
Technology Solutions
Company index
ABB Lummus Global GTC Technology Inc.
Air Products and Chemicals, Inc. Haldor Topsoe
Axens Kobe Steel Ltd.
Bechtel Linde AG
Belco Technologies Corp. Lurgi
CB&I Merichem Chemicals & Reﬁnery Services LLC
CDTECH Process Dynamics, Inc.
Chevron Lummus Global LLC. Reﬁning Hydrocarbon Technology LLC
ConocoPhillips Shaw Stone &Webster
Davy Process Technology Shell Global Solutions International BV
DuPont Technip
ExxonMobil Engineering & Research Uhde GmbH
Foster Wheeler UOP LLC
Gas Technology Products
Genoil Inc.
Goar, Allison & Associates

Sponsored by:
Technology Solutions
ABB Lummus Global
Alkylation
Coking
Fluid catalytic cracking
Hydrotreating
Hydrotreating-aromatic saturation
Air Products and Chemicals, Inc.
Hydrogen-recovery
Olefins recovery
Axens
Alkylation-feed preparation
Benzene reduction
Catalytic reforming
Ethers
Gasoline desulfurization, ultra deep
Hydroconversion-VGO & DAO
Hydrocracking
Hydrocracking-residue
Hydrotreating diesel
Hydrotreating-resid
Isomerization
Lube oil refining, spent
Oligomerization-C3/C4 cuts
Oligomerization-polynaphtha
Spent lube oil re-refining
Bechtel
Dewaxing
Dewaxing/wax deoiling
Lube extraction
Lube extraction
Lube hydrotreating
Lube hydrotreating
Wax hydrotreating
Belco Technologies Corp.
NOx reduction, low-temperature
SO2 removal, regenerative
Wet Scrubbing system, EDV
CB&I
Catalytic reforming
Crude topping units
Hydrogen-steam reforming
Hydrotreating
CDTECH
Alkylation, catalytic
Hydrogenation
Hydrotreating
Isomerization
Chevron Lummus Global LLC.
Dewaxing
Hydrocracking (ISOCRACKING)
Hydrocracking (LC-FINING)
Hydrofinishing
Hydrotreating (ISOTREATING)
Hydrotreating-RDS/VRDS/UFR/OCR
Processes:
Alkylation
Coking Technology Solutions
Gasoline desulfurization Technology Solutions, a division of ConocoPhillips, is a premier provider of technology solutions for
the vehicles of today and the oilfields and energy systems of tomorrow. Backed by modern research facilities
Isomerization and a strong tradition of innovation, we develop, commercialize and license technologies that help oil and
gas producers, refiners and manufacturers reach their business and operational. From enhanced production
methods, to gasoline sulfur removal processes to valuable catalysts that enhance fuel cell operation, Tech-
nology Solutions prepares producers, refiners and consumers alike for a cleaner, more beneficial future.
Strengths of Our Business
• Focused efforts on developing and commercializing technologies that enable refiners to economically
produce clean fuels and upgrade hydrocarbons into higher value products
• Strategic alignment with both Upstream and Downstream energy segments to effectively capitalize on
extensive R&D, commercial and operational expertise
• Strong relationship-building and problem-solving abilities
• Customer inter-facing and advocacy
Industries Served
Technology Solutions supports both Upstream and Downstream energy segments, including:
• Carbon and petroleum coke
• Gasification
• Sulfur chemistry
• Hydrocarbon processing and upgrading
• Upstream technologies
• Enhanced recovery
Corporate Overview
ConocoPhillips (NYSE:COP) is an international, integrated energy company. It is the third-largest integrat-
ed energy company in the United States, based on market capitalization, and oil and gas proved reserves
and production; and the second largest refiner in the United States. Worldwide, of non government-con-
trolled companies, ConocoPhillips has the fifth-largest total of proved reserves and is the fourth-largest
refiner. Headquartered in Houston, Texas, ConocoPhillips operates in more than 40 countries. As of March
31, 2006, the company had approximately 38,000 employees worldwide and assets of $160 billion.

For More Information: ConocoPhillips Technology Solutions

Email: TechnologySolutions@conocophillips.com
Web: www.COPtechnologysolutions.com
Davy Process Technology
Prereforming with feed ultra purification
DuPont
Alkylation
ExxonMobil Engineering & Research
Alkylation-sulfuric acid
Catalytic dewaxing
Coking, fluid
Coking,flexi
Fluid catalytic cracking
Gas treating-H2S removal
Gasoline desulfurization, ultra deep
Gasoline desulfurization, ultra deep
Hydrocracking
Hydroprocessing, ULSD
Lube treating
NOx abatement
Pressure swing adsorption-rapid cycle
Wet gas scrubbing
Processes:
Coking
Crude distillation
Foster Wheeler is a global engineering and construction contractor and power equipment supplier, with
Deasphalting a reputation for delivering high-quality, technically-advanced, reliable facilities and equipment on time, on
budget and with a world-class safety record.
Hydrogen-steam reforming
Our Engineering & Construction Group designs and constructs leading-edge processing facilities for
Visbreaking the upstream oil & gas, LNG & gas-to-liquids, refining, chemicals & petrochemicals, power, environmental,
pharmaceuticals, biotechnology & healthcare industries. Foster Wheeler is a market leader in heavy oil
upgrading technologies, offering world-leading technology in delayed coking, solvent deasphalting, and

Technical articles: visbreaking, and providing cost-effective solutions for the reﬁning industry.

Services:
• Integrated hydrogen solutions: • Market studies
• Master planning
Combining hydrogen recovery • Feasibility studies
and optimized steam • Concept screening
• Environmental engineering
• Upgrade reﬁnery residuals into • Front-end design (FEED)
value-added products • Project management (PMC)
• Engineering (E)
• Optimize turnaround projects • Procurement (P)
• Drivers for additional delayed • Construction (C) & construction management (Cm)
• Commissioning & start-up
coking capacity in the reﬁning • Validation
industry • Plant operations & maintenance
• Training
• When solvent deasphalting is
the most appropriate technology Our Global Power Group, world-leading experts in combustion technology, designs, manufactures and
for upgrading residue erects steam generating and auxiliary equipment for power stations and industrial markets worldwide, and
also provides a range of after-market services.

Email: ann_hooper@fwhou.fwc.com
Web: www.fosterwheeler.com
Gas Technology Products
H2S removal
H2S removal
H2S removal
Genoil Inc.
Hydrotreating-residue
Goar, Allison & Associates
Sulfur processing
Sulfur recovery
GTC Technology Inc.
Aromatics
Aromatics recovery
Desulfurization
Paraxylene
Haldor Topsoe
Diesel-ultra-low-sulfur diesel (ULSD)
Diesel-upgrading
Flue gas denitrification
Flue gas desulfurization-SNOX
Fluid catalytic cracking-pretreatment
H2S and SWS gas conversion
Hydrocracking
Hydrodearmatization
Hydrogen-HTCT and HTCR twin plants
Hydrogen-HTER-p
Hydrogen-methanol-to-shift
Hydrogen-steam-methane reforming (SMR)
Hydrotreating
Sour gas treatment
Spent acid regneration
Kobe Steel Ltd.
Hydrocracking
Linde AG
O2 enrichment for Claus units
O2 enrichment for FCC units
Lurgi
Biodiesel
Merichem Chemicals & Refinery Services LLC
Treating jet fuel/kerosine
Treating-gases
Treating-gasoline and LPG
Treating-gasoline desulfurization, ultra deep
Treating-gasoline sweetening
Treating-kerosine and heavy naphtha sweetening
Treating-phenolic caustic
Treating-propane
Treating-reformer products
Treating-spent caustic deep neutralization
Process Dynamics, Inc.
Hydrotreating
Hydrotreating-lube and wax
Lube and wax processing
Refining Hydrocarbon Technology LLC
Alkylation
Isooctene/isooctane
Olefin etherfication
Shaw Stone & Webster
Deep catalytic cracking
Fluid catalytic cracking
Refinery offgas-purification and olefins recovery
Resid catalytic cracking
Shell Global Solutions International BV
Crude distillation
Deep thermal conversion
Fluid catalytic cracking
Gasification
Hydrocracking
Hydroprocessing, residue
Thermal gasoil
Visbreaking
Technip
Crude distillation
Hydrogen
Uhde GmbH
Aromatics extractive distillation
Ethers-ETBE
Ethers-MTBE
Hydrofinishing/hydrotreating
Hydrogen
Lube treating
Olefins-butenes extractive distillation
Olefins-dehydrogenation of light parraffins to olefins
Slack wax deoiling
Vacuum distillation
White oil and wax hydrotreating
Processes:
Alkylation (2)
Alkylation-HF For more than 90 years, UOP LLC, a Honeywell company, has been a leader in developing and com-
mercializing technology for license to the oil refining, petrochemical and gas processing industries. Starting
Catalytic reforming with its first breakthrough technology, UOP has contributed processes and technology that have led to ad-
vances in such diverse industries as motor fuels, plastics, detergents, synthetic fibers and food preservatives.
Fluid catalytic cracking UOP is the largest process licensing organization in the world, providing more than 50 licensed processes
Hydrocracking for the hydrocarbon processing industries and holding more than 2,500 active patents.
UOP offices are in Des Plaines, Illinois, USA (a northwest suburb of Chicago). The company employs
Hydrotreating (2) nearly 3,000 people in its facilities in the United States, Europe and Asia.
The petroleum refining industry is the largest market for UOP technology, products and services. UOP
Hydrotreating/desulfurization processes are used throughout the industry to produce clean-burning, high-performance fuels from a vari-
ety of hydrocarbon products. For example, for 60 years our Platforming process has been used to upgrade
Isomerization (3) low-octane naphtha to high-octane unleaded gasoline, a higher performance fuel. Other technologies con-
vert mercaptans to innocuous disulfides, remove sulfur from fuel, and recover high-purity hydrogen from
Mercaptan removal impure gas streams.
Technologies developed by UOP are almost entirely responsible for providing the fundamental raw
Treating-pressure swing materials – benzene, toluene and xylene (BTX) – of the aromatics-based petrochemicals industry. These
products form the basis of such familiar products as synthetic rubber, polyester fibers, polystyrene foam,
adsorption glues and pharmaceuticals. UOP technologies produce such olefins as ethylene and propylene, used in a
range of products from contact lenses to food packaging. UOP has been active in the development of syn-
thetic detergent chemicals since 1947, and today almost half of the world’s soft (biodegradable) detergents
Technical articles: are produced through UOP-developed processes.
UOP’s gas processing technologies are used for separating, drying and treating gases produced from oil
• Concepts for an overall refinery and gas wells and atmospheric gases.
energy solution through novel in- UOP is the world’s leading producer of synthetic molecular sieve adsorbents used in purifying natural gas,
separating paraffins and drying air through cryogenic separation. Molecular sieves also are used in insulat-
tegration of FCC flue gas power ing glass, refrigeration systems, air brake systems, automotive mufflers and deodorizing products.
recovery UOP provides engineering designs for its processes, and produces key mechanical equipment for some of
its processes. It also offers project management, cost estimation, procurement and facility-design services.
• Changing refinery configura- UOP’s staff of engineers provides customers with a wide range of services, including start-up assistance,
tion for heavy and synthetic operating technical services such as process monitoring and optimization, training of customer personnel,
crude processing catalyst and product testing, equipment inspection, and project management.
For more information:
jennifer.wilson@uop.com
Alkylation
Application: The AlkyClean process converts light olefins into alkylate
by reacting the olefins with isobutane over a true solid acid catalyst.
AlkyClean’s unique catalyst, reactor design and process scheme allow ��
operation at low external isobutane-to-olefin ratios while maintaining
excellent product quality. ��
��
��
Products: Alkylate is a high-octane, low-Rvp gasoline component used ��
��
��
for blending in all grades of gasoline. ��
��
Description: The light olefin feed is combined with the isobutane make-
��
up and recycle and sent to the alkylation reactors which convert the ��
��
olefins into alkylate using a solid acid catalyst (1). The AlkyClean process ��
uses a true solid acid catalyst to produce alkylate, eliminating the safety
and environmental hazards associated with liquid acid technologies. Si-
multaneously, reactors are undergoing a mild liquid-phase regeneration
using isobutane and hydrogen and, periodically, a reactor undergoes a
higher temperature vapor phase hydrogen strip (2). The reactor and mild
regeneration effluent is sent to the product-fractionation section, which
produces n-butane and alkylate products, while also recycling isobutane
and recovering hydrogen used in regeneration for reuse in other refinery Installation: Demonstration unit at Neste Oil’s Porvoo, Finland Refinery.
hydroprocessing units (3). The AlkyClean process does not produce any
acid soluble oils (ASO) or require post treatment of the reactor effluent or Reference: “The AlkyClean process: New technology eliminates liquid
final products. acids,” NPRA 2006 Annual Meeting, March 19–21, 2006.
D’Amico, V., J. Gieseman, E. von Broekhoven, E. van Rooijen and
Product: The C5+ alkylate has a RON of 93–98 depending on processing H. Nousiainen, “Consider new methods to debottleneck clean alkylate
conditions and feed composition. production,” Hydrocarbon Processing, February 2006, pp. 65–70.
Economics: Licensor: ABB Lummus Global, Albemarle Catalysts and Neste Oil.
Investment (2006 USGC basis 10,000-bpsd unit) $/bpsd 4,200
Operating cost, $/gal 0.08
Alkylation
Application: Convert propylene, butylenes, amylenes and isobutane to
the highest quality motor fuel using ReVAP (Reduce Volatility Alkylation
��
Process) alkylation.
�
Products: An ultra-low-sulfur, high-octane and low-Rvp blending stock
��
for motor and aviation fuels.
��

Description: Dry liquid feed containing oleﬁns and isobutane is charged ��
�
to a combined reactor-settler (1). The reactor uses the principle of dif-
ferential gravity head to effect catalyst circulation through a cooler pri- ��
or to contacting highly dispersed hydrocarbon in the reactor pipe. The
��
hydrocarbon phase that is produced in the settler is fed to the main
��
fractionator (2), which separates LPG-quality propane, isobutane recycle, ��

n-butane and alkylate products. A small amount of dissolved catalyst is

removed from the propane product by a small stripper tower (3). Major
process features are:
• Gravity catalyst circulation (no catalyst circulation pumps re-
quired)
• Low catalyst consumption Yields: Feed type
Butylene Propylene-butylene mix
• Low operating cost
Alkylate product
• Superior alkylate qualities from propylene, isobutylene and amyl-
Gravity, API 70.1 71.1
ene feedstocks
Rvp, psi 6–7 6–7
• Onsite catalyst regeneration
ASTM 10%, °F 185 170
• Environmentally responsible (very low emissions/waste)
ASTM 90%, °F 236 253
• Between 60% and 90% reduction in airborne catalyst release over
RONC 96.0 93.5
traditional catalysts
Per bbl olefin converted
• Can be installed in all licensors’ HF alkylation units.
i-Butane consumed, bbl 1.139 1.175
With the proposed reduction of MTBE in gasoline, ReVAP offers sig-
Alkylate produced, bbl 1.780 1.755
niﬁcant advantages over sending the isobutylene to a sulfuric-acid-al-
kylation unit or a dimerization plant. ReVAP alkylation produces higher Installation: 147 worldwide licenses.
octane, lower Rvp and lower endpoint product than a sulfuric-acid-alkyl-
ation unit and nearly twice as many octane barrels as can be produced Licensor: ConocoPhillips.
from a dimerization unit.
Alkylation
��
Application: To combine propylene, butylenes and amylenes with isobutane ��
in the presence of strong sulfuric acid to produce high-octane branched �
chain hydrocarbons using the Efﬂuent Refrigeration Alkylation process.
� ��
Products: Branched chain hydrocarbons for use in high-octane motor � ��
�
fuel and aviation gasoline. �

��
Description: Plants are designed to process a mixture of propylene, ��
butylenes and amylenes. Oleﬁns and isobutane-rich streams along with ��
a recycle stream of H2SO4 are charged to the STRATCO Contactor reac- ��

tor (1). The liquid contents of the Contactor reactor are circulated at high
velocities and an extremely large amount of interfacial area is exposed ��
between the reacting hydrocarbons and the acid catalyst from the acid ��

settler (2). The entire volume of the liquid in the Contactor reactor is main-
tained at a uniform temperature, less than 1°F between any two points
within the reaction mass. Contactor reactor products pass through a ﬂash
drum (3) and deisobutanizer (4). The refrigeration section consists of a
compressor (5) and depropanizer (6).
The overhead from the deisobutanizer (4) and efﬂuent refrigerant Utilities, typical per bbl alkylate:
recycle (6) constitutes the total isobutane recycle to the reaction zone. Electricity, kWh 13.5
This total quantity of isobutane and all other hydrocarbons is maintained Steam, 150 psig, lb 180
in the liquid phase throughout the Contactor reactor, thereby serving to Water, cooling (20°F rise), 103 gal 1.85
promote the alkylation reaction. Onsite acid regeneration technology is Acid, lb 15
also available. Caustic, lb 0.1

Product quality: The total debutanized alkylate has RON of 92 to 96 Installation: Over 600,000 bpsd installed capacity.
clear and MON of 90 to 94 clear. When processing straight butylenes,
Reference: Hydrocarbon Processing, Vol. 64, No. 9, September 1985,
the debutanized total alkylate has RON as high as 98 clear. Endpoint of
pp. 67–71.
the total alkylate from straight butylene feeds is less than 390°F, and less
than 420°F for mixed feeds containing amylenes in most cases. Licensor: DuPont.
Economics (basis: butylene feed):
Investment (basis: 10,000-bpsd unit), $ per bpsd 4,500
Alkylation
Application: The RHT-Alkylation process is an improved method to react ��
C3– C5 olefins with isobutane using the classical sulfuric acid alkylation
��
process. This process uses a unique mixing device — eductor(s) — that
��
provides low-temperature (25 – 30°F) operations at isothermal condi- ��
�
��
�
��
tions. This eductor mixing device is more cost-effective than other de-
vices being used or proposed. It is maintenance free and does not re- ��
��
quire replacement every two to three years. This mixing device can be a
�
retrofit replacement for existing contactors. In addition, the auto refrig- ��
��
eration vapor can be condensed by enhancing pressure and then easily ��
��
��
absorbed in hydrocarbon liquid, without revamping the compressor. ��
��
��
Description: In the RHT-Alkylation, C3– C5 feed from FCC or any other ��
��
source including steam cracker, etc., with isobutane make-up, recycle
isobutene, and recovered hydrocarbons from the depropanizer bottom ��
and refrigeration vapors are collected in a surge drum — the C4 system
(5). The mixture is pumped to the reactor (1) to the eductor suction port.
The motive fluid is sent to the eductor nozzle from the bottom of reac-
tor, which is essentially sulfuric acid, through pumps to mix the reactants
with the sulfuric-acid catalyst.
The mixing is vigorous to move the reaction to completion. The The hydrocarbon is sent to distillation column(s) (7), to separate alkyl-
makeup acid and acid-soluble oil (ASO) is removed from the pump dis- ate product and isobutane, which is recycled. The butane is sent to offsites
charge. The process has provisions to install a static mixer at the pump or can be converted back to isobutane for processing units requirements.
discharge. Some feed can be injected here to provide higher OSV, which The auto refrigeration occurs in the reactor at temperatures 25–30°F. The
is required for C3 alkylation. Reactor effluent is withdrawn from the isothermal condition lowers acid consumption and yields higher octane
reactor as a side draw and is sent to acid/ hydrocarbon coalescer (2) product due to improved selectivity of 2,4,4 trimethylpentane.
where most of the acid is removed and recycled to the reactor (1). The The auto-refrigeration vapor is compressed (or first enhanced the
coalescers are being used by conventional process to reduce the acid in pressure by the ejector and then absorbed in a heavy liquid — alkylate,
the hydrocarbon phase to 7–15 wppm. The enhanced coalescer design which provides a low-cost option) and then condensed. Some liquid is
RHT can reduce the sulfuric acid content in the hydrocarbon phase to sent to depropanizer (6); propane and light ends are removed. The bot-
negligible levels (below <1 wppm). toms are recycled to C4 system and sent to the reactor.
After the coalescer, the hydrocarbon phase is heated and flashed The major advances of RHT process are threefold: eductor mixing
increasing the alkylate concentration in the hydrocarbon, which is sent device, advance coalescer system to remove acid from hydrocarbon (dry
through the finishing coalescer where essentially all of the remaining system), and C4 autorefrigeration vapors recovery by absorption, mak-
acid is removed. ing compressor redundant.
Continued 
Alkylation, continued Commercial units: Technology is ready for commercialization.
References:
Economics: For a US Gulf Coast unit 1Q 2006 with a capacity of 10,000 US patent 5,095168.
bpd alkylate unit US Patent 4,130593.
CAPEX ISBL, MM USD 31.2 Kranz, K., “Alkylation Chemistry,” Stratco, Inc., 2001.
Utilities ISBL costs, USD/ bbl alkylate 3,000 Branzaru, J., “Introduction to Sulfuric Acid Alkylation,” Stratco, Inc.,
Power, kWh 4,050* 2001.
Water, cooling, m3/ h 1,950 Nelson, Handbook of Refining.
Steam, kg / h 25,600 Meyers, R. A., Handbook of Refining, McGraw Hill, New York,
1997.
* Power could be less for absorption application
FCC Feed (about 15% isobutelene in C4 mixed stream) Licensor: Refining Hydrocarbon Technologies LLC.
Product properties: Octane (R+M) / 2:94.8 – 95.4
Alkylation
Application: The Alkad process is used with HF alkylation technology to ��
��
reduce aerosol formation in the event of an HF release, while maintain-
��
ing unit operability and product quality. The Alkad process is a passive
mitigation system that will reduce aerosol from any leak that occurs
while additive is in the system. ��
��
��
Description: The additive stripper sends acid, water and light-acid sol- ��
uble oils overhead and on to the acid regenerator. Heavy acid soluble ��
oils and the concentrated HF-additive complex are sent to the additive
��
stripper bottoms separator. From this separator the polymer is sent to ��
��
neutralization, and the HF-additive complex is recycled to the reactor
section. The acid regenerator removes water and light-acid soluble oils
from the additive stripper overhead stream. The water is in the form of ��
a constant boiling mixture (CBM) of water and HF. ��
��
There is no expected increase in the need for operator manpower.
Maintenance requirements are similar to equipment currently in stan-
dard operation in an HF alkylation unit in similar service.

Experience: ChevronTexaco, the co-developer of the Alkad process, in-

stalled facilities to use this technology in the HF Alkylation unit at their
former El Dorado, Kansas, reﬁnery. This unit began initial operations in
1994.

Installation: One unit is under construction.

Licensor: UOP LLC and Chevron Corp.
Alkylation
Application: UOP’s Indirect Alkylation (InAlk) process uses solid catalysts ��
to convert light olefins (mainly C4 but also C3 and C5) to alkylate. ��
��
��
Description: The InAlk process makes premium alkylate using a combi-
nation of commercially proven technologies. Iso-butene reacts with itself
or with other C3– C5 olefins via polymerization. The resulting mixture of
�
higher molecular weight iso-olefins may then be hydrogenated to form � �
� �
a high-octane paraffinic gasoline blendstock that is similar to alkylate,
but usually higher in octane, or it may be left as an olefinic high-octane
gasoline blending component.
Either resin or solid phosphoric acid (SPA) catalysts are used to po-
lymerize the olefins. Resin catalyst primarily converts iso-butene. SPA
catalyst also converts n-butenes. The saturation section uses either a
��
base-metal or noble-metal catalyst. ��

Feed: A wide variety of feeds can be processed in the InAlk process.

Typical feeds include FCC-derived light olefins, steam-cracker olefins
and iC4 dehydrogenation olefins.

Installation: The InAlk process is an extension of UOP’s catalytic con-

densation and olefin saturation technologies. UOP has licensed and de-
signed more than 400 catalytic condensation units for the production of
polygasoline and petrochemical olefins and more than 200 hydrogena-
tion units of various types. Currently five InAlk units are in operation.

Licensor: UOP LLC.

Alkylation, catalytic ��
Application: CDAlky process is an advanced sulfuric acid-catalyzed al-
kylation process that reacts light olefin streams from refinery sources, ��
such as fluid catalytic cracking (FCC) units or from steam-cracking units,
with iso-paraffins to produce motor fuel alkylate.

Description: The patented CDAlky process is an advanced sulfuric acid-

��
catalyzed alkylation process for the production of motor fuel alkylate. The
��
process flow diagram shows the basic configuration to process a mixed ��
C4-olefin feed and produce a bright, clear, high-quality motor fuel alkyl- ��
ate, without the need for water/caustic washes or bauxite treatment. ��
This process yields a higher-quality product while consuming sig- ��
nificantly less acid than conventional technologies. The flow scheme is ��
also less complex than conventional designs, which reduces capital and
��
operating costs.
Conventional sulfuric-acid alkylation units use mechanical mixing ��
in their contactors to achieve the required contact between acid and
hydrocarbon phases, and are characterized by high acid consumption.
In addition, conventional technologies are unable to take the full ben-
efit of operating at lower temperature, which substantially improves
alkylate quality and lowers acid consumption. • Lower maintenance—no mechanical agitator or complex seals
CDTECH has developed a novel contactor that operates at lower • Less corrosion due to dry system
temperatures and substantially reduced acid consumption—50%+. The • No caustic waste stream
CDAlky process uses conventional product fractionation, which can con- Installation: Consistent with time-tested methodology for developing
sist of a single column or two columns. This process has been designed new processes, CDTECH has been operating a 2-bpd pilot plant in this
to make it possible to reuse equipment from idled facilities. novel mode of operation for an extended time period without the pen-
The benefits of the CDAlky process include: alties associated with conventional technologies.
• Lower acid consumption
• Lower utilities Licensor: CDTECH.
• Reduced operating cost
• Reduced environmental exposure
• Higher octane product
• Lower CAPEX—simpler flowsheet with fewer pieces of equipment
• Highly flexible operation range—maximum absolute product oc-
tane or maximum octane barrels
Alkylation—feed preparation
Application: Upgrades alkylation plant feeds with Alkyfining process.
��
Description: Diolefins and acetylenes in the C4 (or C3– C4) feed react se-
lectively with hydrogen in the liquid-phase, fixed-bed reactor under mild
temperature and pressure conditions. Butadiene and, if C3s are present,
methylacetylene and propadiene are converted to olefins.
The high isomerization activity of the catalyst transforms 1-butene
into cis- and trans-2-butenes, which affords higher octane-barrel pro- ��

duction.
Good hydrogen distribution and reactor design eliminate channeling
��
while enabling high turndown ratios. Butene yields are maximized, hy- ��
drogen is completely consumed and, essentially, no gaseous byproducts
or heavier compounds are formed. Additional savings are possible when
pure hydrogen is available, eliminating the need for a stabilizer. The pro-
cess integrates easily with the C3/C4 splitter.
Alkyfining performance and impact on HF alkylation product:
The results of an Alkyﬁning unit treating an FCC C4 HF alkylation Annual savings for a 10,000-bpsd alkylation unit:
unit feed containing 0.8% 1,3-butadiene are: HF unit, US$ 4.1 million
Butadiene in alkylate, ppm < 10 H2SO4 unit, US$ 5.5 million
1-butene isomerization, % 70
Butenes yield, % 100.5
Installation: Over 90 units are operating with a total installed capacity
of 800,000 bpsd.
RON increase in alkylate 2
MON increase in alkylate 1 Licensor: Axens.
Alkylate end point reduction, °C –20
The increases in MON, RON and butenes yield are reﬂected in a
substantial octane-barrel increase while the lower alkylate end point re-
duces ASO production and HF consumption.

Economics:
Investment:
New unit ISBL cost:
For an HF unit, $/bpsd 430
For an H2SO4 unit, $/bpsd 210
Alkylation—HF
Application: HF Alkylation improves gasoline quality by adding clean-
burning, mid-boiling-range isoparaffins and reducing gasoline pool va-
por pressure and olefin content by conversion of C3– C5 olefin compo- ��
��
nents to alkylate.
Description: The alkylation reaction catalytically combines C3– C5 olefins ��
with isobutane to produce motor-fuel alkylate. Alkylation takes place in ��
��
the presence of HF catalyst under conditions selected to maximize alkyl-
ate yield and quality.
The reactor system is carefully designed to ensure efficient contact- ��
��
��
ing and mixing of hydrocarbon feed with the acid catalyst. Efficient heat ��
��
transfer conserves cooling water supply. Acid inventory in the reactor
system is minimized by combining high heat-transfer rates and lower ��
total acid circulation.
Acid regeneration occurs in the acid regenerator or via a patented
internal-acid-regeneration method. Internal regeneration allows the
refiner to shutdown the acid regenerator, thereby realizing a utility
savings as well as reducing acid consumption and eliminating polymer
disposal.
Feed: Alkylation feedstocks are typically treated to remove sulfur and
water. In cases where MTBE and TAME raffinates are still being pro-
cessed, an oxygenate removal unit (ORU) may be desirable.
Selective hydrogenation of butylene feedstock is recommended to
reduce acid regeneration requirements, catalyst (acid) consumption and
increase alkylate octane by isomerizing 1-butene to 2-butene.
Efficiency: HF Alkylation remains the most economically viable method
for the production of alkylate. The acid consumption rate for HF Alkyla-
tion is less than 1/100th the rate for sulfuric alkylation units. And un-
like sulfuric alkylation units, HF Alkylation does not require refrigeration
equipment to maintain a low reactor temperature.
Installations: Over 20 UOP licensed HF alkylation units are in operation.
Licensor: UOP LLC.
Alkylation, sulfuric acid
Application: Autorefrigerated sulfuric-acid catalyzed process that com- Propane product
bines butylene (and propylene or pentylene if desired) with isobutane
to produce high-octane gasoline components that are particularly at-
tractive in locations that are MON limited. Technology can be installed 3
2
grassroots or retrofit into existing alkylation facilities. Recycle
Refrigerant Butane
isobutane product
Products: A low-sulfur, low-Rvp, highly isoparaffinic, high-octane (espe-
cially MON) gasoline blendstock is produced from this alkylation process.
1
4
Description: Olefin feed and recycled isobutane are introduced into the 5 6
stirred, autorefrigerated reactor (1). Mixers provide intimate contact be-
Olefin feed
tween the reactants and acid catalyst. Highly efficient autorefrigeration
removes heat of reaction heat from the reactor. Hydrocarbons, vaporized START
Alkylate
Makeup
from the reactor to provide cooling, are compressed (2) and returned to Recycle acid isobutane product
the reactor. A depropanizer (3), which is fed by a slipstream from the
refrigeration section, is designed to remove any propane introduced to
the plant with the feeds.
Hydrocarbon products are separated from the acid in the settler
containing proprietary internals (4). In the deisobutanizer (5) isobutane
is recovered and recycled along with makeup isobutane to the reactor. Steam, lb 200
Butane is removed from alkylate in the debutanizer (6) to produce a H2SO4, lb 19
low-Rvp, high-octane alkylate product. A small acid stream containing NaOH, 100%, lb 0.1
acid soluble oil byproducts is removed from the unit and is either regen-
erated on site or sent to an off-site sulfuric acid regeneration facility to
Operating experience: Extensive commercial experience in both
ExxonMobil and licensee refineries, with a total operating capacity of
recover acid strength.
119,000-bpsd at 11 locations worldwide. Unit capacities currently range
Yields: from 2,000 to 30,000 bpd. The license of the world’s largest alkylation
Alkylate yield 1.8 bbl C5+/ bbl butylene feed unit, with a capacity of 83,000 bpd, was recently announced at Reliance
Isobutane required 1.2 bbl / bbl butylene feed Petroleum Limited’s Export Refinery in Jamnagar, India. A revamp has
Alkylate quality 97 RON / 94 MON been completed at ExxonMobil’s Altona, Australia refinery and a new
Rvp, psi 3 unit at TNK-BP’s Ryazan, Russia refinery is scheduled to start-up in mid-
2006. The larger units take advantage of the single reactor/settler trains
Utilities: typical per barrel of alkylate produced: with capacities up to 9,500 bpsd.
Water, cooling, M gal 2
Power, kWH 9 Continued 
Alkylation, sulfuric acid, continued Economic advantages:
• Lower capital investment—Simple reactor/settler configura-
tion, less compression requirements translate into a significant invest-
Technical advantages: ment savings compared to indirect refrigeration systems
• Autorefridgeration is thermodynamically more efficient, allows
• Lower operating costs—Autorefrigeration, lower mixing and
lower reactor temperatures, which favor better product quality, and
compression power requirements translate into lower operating costs
lowers energy usage.
• Better economy of scale —Reactor system is simple and easily
• Staged reactor results in a high average isobutane concentra-
expandable with 9,500 bpsd single train capacities easily achievable.
tion, which favors high product quality.
• Low space velocity results in high product quality and reduced Reference: Lerner, H., “Exxon sulfuric acid alkylation technology,” Hand-
ester formation eliminating corrosion problems in fractionation equip- book of Petroleum Refining Processes, 2nd Ed., R. A. Meyers, Ed., pp.
ment. 1.3–1.14.
• Low reactor operating pressure translates into high reliabil-
ity for the mechanical seals for the mixers, which operate in the vapor Licensor: ExxonMobil Research & Engineering Co.
phase.
Aromatics
��
Application: The GT-TransAlk technology produces benzene and xylenes ��
from toluene and/or heavy aromatics streams. The technology features ��
��
a proprietary catalyst and can accommodate varying ratios of feedstock,
while maintaining high activity and selectivity.
��
Description: The GT-TransAlk technology encompasses three main pro-
��
cessing areas: feed preparation, reactor and product stabilization sec- ��
��
tions. The heavy aromatics stream (usually derived from catalytic refor-
mate) is fed to a C10/C11 splitter. The overhead portion, along with any ��
��
toluene that may be available, is the feed to the transalkylation reactor ��
section. The combined feed is mixed with hydrogen, vaporized, and fed ��
to the reactor. The un-reacted hydrogen is recycled for re-use. The prod-
��
uct stream is stabilized to remove fuel gas and other light components. ��
The process reactor is charged with a proprietary catalyst, which ��
��
exhibits good flexibility to feed stream variations, including 100% C9+
aromatics. Depending on the feed composition, the xylene yield can
vary from 27 to 35% and C9 conversion from 53 to 67%.
Process advantages include:
• Simple, low cost fixed-bed reactor design
• Selective toward xylene production, with high toluene/C9 conver-
sion rates
• Physically stable catalyst
• Flexible to handle up to 100% C9+ components in feed
• Flexible to handle benzene recycle to increase xylene yields
• Moderate operating parameters; catalyst can be used as replace-
ment to other transalkylation units, or in grassroots designs
• Decreased hydrogen consumption due to low cracking rates
• Efficient heat integration scheme, reduces energy consumption.

Licensor: GTC Technology Inc.

Aromatics extractive distillation
��
��
Application: Recovery of high-purity aromatics from reformate, pyrolysis ��
gasoline or coke oven light oil using extractive distillation. ��

Description: In Uhde’s proprietary extractive distillation (ED) Morphylane

process, a single-compound solvent, N-Formylmorpholine (NFM), alters the
��
vapor pressure of the components being separated. The vapor pressure of ��
the aromatics is lowered more than that of the less soluble nonaromatics.
Nonaromatics vapors leave the top of the ED column with some solv- ��
ent, which is recovered in a small column that can either be mounted on ��
��
the main column or installed separately.
Bottom product of the ED column is fed to the stripper to separate
pure aromatics from the solvent. After intensive heat exchange, the lean
solvent is recycled to the ED column. NFM perfectly satisfies the neces- ��
��
sary solvent properties needed for this process including high selectivity,
thermal stability and a suitable boiling point.
Uhde’s new single-column morphylane extractive distillation process
uses a single-column configuration, which integrates the ED column and
the stripper column of the conventional design. It represents a superior
Reformate feedstock with low-aromatics content (20 wt%):
process option in terms of investment and operating cost.
Benzene
Economics: Quality
Benzene 10 wt ppm NA*
Pygas feedstock:
Production yield Benzene Benzene/toluene Consumption
Steam 320 kg/t ED feed
Benzene 99.95% 99.95%
Toluene – 99.98% *Maximum content of nonaromatics **Including benzene/toluene splitter
Quality
Benzene 30 wt ppm NA* 80 wt ppm NA* Installation: More than 55 Morphylane plants (total capacity of
Toluene – 600 wt ppm NA* more than 6 MMtpy). The first single-column Morphylane unit went
Consumption onstream in 2004.
Steam 475 kg/t ED feed 680 kg/t ED feed**
References: Diehl, T., B. Kolbe and H. Gehrke, “Uhde Morphylane Ex-
tractive Distillation—Where do we stand?” ERTC Petrochemical Confer-
ence, October 3–5, 2005, Prague.

Continued 
Aromatics extractive distillation, continued
Emmrich, G., U. Ranke and H. Gehrke, “Working with an extractive dis-
tillation process,” Petroleum Technology Quarterly, Summer 2001, p. 125.

Licensor: Uhde GmbH.

Aromatics recovery
Application: GT-BTX is an aromatics recovery process. The technology
uses extractive distillation to remove benzene, toluene and xylene (BTX)
from reﬁnery or petrochemical aromatics streams such as catalytic re-
formate or pyrolysis gasoline. The process is superior to conventional
liquid-liquid and other extraction processes in terms of lower capital and ��
��
operating costs, simplicity of operation, range of feedstock and solvent ��
��
��
performance. Flexibility of design allows its use for grassroots aromatics ��
��
��
recovery units, debottlenecking or expansion of conventional extraction ��
systems. �

Description: The technology has several advantages:

• Less equipment required, thus, significantly lower capital cost
compared to conventional liquid-liquid extraction systems ��
• Energy integration reduces operating costs
• Higher product purity and aromatic recovery
• Recovers aromatics from full-range BTX feedstock without pre-
fractionation
• Distillation-based operation provides better control and simplified
operation
• Proprietary formulation of commercially available solvents exhibits reflux to wash out any entrained solvent. The balance of the overhead
high selectivity and capacity stream is the raffinate product, requiring no further treatment.
• Low solvent circulation rates Rich solvent from the bottom of the EDC is routed to the solvent-re-
• Insignificant fouling due to elimination of liquid-liquid contactors covery column (SRC), where the aromatics are stripped overhead. Strip-
• Fewer hydrocarbon emission sources for environmental benefits ping steam from a closed-loop water circuit facilitates hydrocarbon re-
• Flexibility of design options for grassroots plants or expansion of moval. The SRC is operated under a vacuum to reduce the boiling point
existing liquid-liquid extraction units at the base of the column. Lean solvent from the bottom of the SRC
• Design avoids contamination of downsream products by objec- is passed through heat exchange before returning to the EDC. A small
tionable solvent carryover. portion of the lean circulating solvent is processed in a solvent-regenera-
Hydrocarbon feed is preheated with hot circulating solvent and fed tion step to remove heavy decomposition products.
at a midpoint into the extractive distillation column (EDC). Lean solvent The SRC overhead mixed aromatics product is routed to the purifi-
is fed at an upper point to selectively extract the aromatics into the col- cation section, where it is fractionated to produce chemical-grade ben-
umn bottoms in a vapor/liquid distillation operation. The nonaromatic zene, toluene and xylenes.
hydrocarbons exit the top of the column and pass through a condenser.
A portion of the overhead stream is returned to the top of the column as Continued 
Aromatics recovery, continued
Economics: Estimated installed cost for a 15,000-bpd GT-BTX extraction
unit processing BT-reformate feedstock is $12 million (US Gulf Coast
2004 basis).

Installations: Fourteen licenses placed.

Licensor: GTC Technology Inc.
Benzene reduction
Application: Benzene reduction from reformate, with the Benfree pro-
cess, using integrated reactive distillation. ��
��
Description: Full-range reformate from either a semiregenerative or CCR
reformer is fed to the reformate splitter column, shown above. The split- ��
ter operates as a dehexanizer lifting C6 and lower-boiling components
to the overhead section of the column. Benzene is lifted with the light ��
ends, but toluene is not. Since benzene forms azeotropic mixtures with
some C7 paraffin isomers, these fractions are also entrained with the
��
light fraction.
Above the feed injection tray, a benzene-rich light fraction is withdrawn
and pumped to the hydrogenation reactor outside the column. A pump
enables the reactor to operate at higher pressure than the column, thus ��
ensuring increased solubility of hydrogen in the feed.
A slightly higher-than-chemical stoichiometric ratio of hydrogen to ben-
zene is added to the feed to ensure that the benzene content of the
resulting gasoline pool is below mandated levels, i.e., below 1.0 vol%
for many major markets. The low hydrogen flow minimizes losses of
gasoline product in the offgas of the column. Benzene conversion to Economics:
cyclohexane can easily be increased if even lower benzene content is Investment, New unit ISBL cost, $/bpsd: 300
desired. The reactor effluent, essentially benzene-free, is returned to the Combined utilities, $/bbl 0.17
column. Hydrogen Stoichiometric to benzene
Catalyst, $/bbl 0.01
The absence of benzene disrupts the benzene-iso-C7 azeotropes, there-
by ensuring that the latter components leave with the bottoms fraction Installation: Twenty-eight benzene reduction units have been licensed.
of the column. This is particularly advantageous when the light refor-
mate is destined to be isomerized, because iso-C7 paraffins tend to be Licensor: Axens.
cracked to C3 and C4 components, thus leading to a loss of gasoline
production.
Biodiesel
Application: Consumption of primary energy has risen substantially in
recent years, and greenhouse gases (GHG) emissions have increased by
��
a substantial amount. To counter this trend, there is a global strong em-
phasis on regenerative energy such as biofuels to effectively reduce or ��
avoid such emissions.

Description: The Lurgi biodiesel process is centered on the transesteriﬁ- ��

��
cation of different raw materials to methyl ester using methanol in the ��
��
presence of a catalyst. In principle, most edible oils and fats — both veg- ��
etable and animal sources— can be transesterified if suitably prepared. ��
Transesterification is based on the chemical reaction of triglycerides ��
with methanol to methyl ester and glycerine in the presence of an alka- ��
��
line catalyst. The reaction occurs in two mixer-settler units. The actual
conversion occurs in the mixers. The separation of methyl ester as the
light phase and glycerine water as the heavy phase occurs in the set-
tlers due to the insolubility of both products. Byproduct components
are removed from the methyl ester in the downstream washing stage,
which operates in a counter-current mode. After a final drying step un-
der vacuum, the biodiesel is ready for use.
Any residual methanol contained in the glycerine water is removed Methanol, kg 96
in a rectification column. In this unit operation, the methanol has a puri- Catalyst (Na-Methylate 100%), kg 5
ty, which is suitable for recycling back to process. For further refinement Hydrochloric Acid (37%), kg 10
of the glycerine water, optional steps are available such as chemical Caustic soda (50%), kg 1.5
treatment, evaporation, distillation and bleaching to either deliver crude Nitrogen, Nm3 1
glycerine at approximately 80% concentration or pharmaceutical-grade Installation: Lurgi has been building biodiesel plants for 20 years. Only in
glycerine at >99.7% purity. the last five years, Lurgi has contracted more than 40 plants for the pro-
Economics: The (approximate) consumption figures—without glycerine duction of biodiesel with capacities ranging from 30,000 to 200,000 tpy.
distillation and bleaching—stated below are valid for the production of Licensors: Lurgi AG.
one ton of rapeseed methyl ester at continuous operation and nominal
capacity.
Steam, kg 320
Water, cooling water (t = 10°C), m3 25
Electrical energy, kWh 12
Catalytic reforming
Application: Upgrade various types of naphtha to produce high-octane
reformate, BTX and LPG.

Description: Two different designs are offered. One design is conventional

where the catalyst is regenerated in place at the end of each cycle. Oper- ��
ating normally in a pressure range of 12 to 25 kg /cm2 (170 to 350 psig) ��

and with low pressure drop in the hydrogen loop, the product is 90 to 100 �
� �
RONC. With its higher selectivity, trimetallic catalysts RG582 and RG682 �
make an excellent catalyst replacement for semi-regenerative reformers.
The second, the advanced Octanizing process, uses continuous cata-
lyst regeneration allowing operating pressures as low as 3.5 kg /cm2 (50
psig). This is made possible by smooth-ﬂowing moving bed reactors (1–3)
which use a highly stable and selective catalyst suitable for continuous
regeneration (4). Main features of Axens’ regenerative technology are: ��

• Side-by-side reactor arrangement, which is very easy to erect and

consequently leads to low investment cost.
• The Regen C2 catalyst regeneration system featuring the dry burn
loop, completely restores the catalyst activity while maintaining its
specific area for more than 600 cycles.
Finally, with the new CR401 (gasoline mode) and AR501 (aromatics Economics:
production) catalysts specifically developed for ultra-low operating pres- Investment: Basis 25,000 bpsd continuous Octanizing unit, battery
sure and the very effective catalyst regeneration system, refiners operat- limits, erected cost, US$ per bpsd 1,800
ing Octanizing or Aromizing processes can obtain the highest hydrogen, Utilities: typical per bbl feed:
C5+ and aromatics yields over the entire catalyst life. Fuel, 10 3 kcal 65
Electricity, kWh 0.96
Yields: Typical for a 90°C to 170°C (176°F to 338°F) cut from light Ara-
Steam, net, HP, kg 12.5
bian feedstock: Conventional Octanizing Water, boiler feed, m3 0.03
Oper. press., kg /cm2 10 –15 <5
Yield, wt% of feed: Installation: Of 130 units licensed, 75 units are designed with continu-
Hydrogen 2.8 3.8 ous regeneration technology capability.
C5+ 83 88
RONC 100 102 Reference: “Octanizing reformer options to optimize existing assets,”
MONC 89 90.5 NPRA Annual Meeting, March 15 –17, 2005, San Francisco.
Continued 
Catalytic reforming, continued
“Fixed Bed Reformer Revamp Solutions for Gasoline Pool Improve-
ment,” Petroleum Technology Quarterly, Summer 2000.
“Increase reformer performance through catalytic solutions,” Sev-
enth ERTC, November 2002, Paris.
“Squeezing the most out of fixed-bed reactors,” Hart Show Special,
NPRA 2003 Annual.

Licensor: Axens.
Catalytic dewaxing
Application: Use the ExxonMobil Selective Catalytic Dewaxing (MSDW) ��
process to make high VI lube base stock. ��
��
Products: High VI / low-aromatics lube base oils (light neutral through ��
��
��
bright stocks). Byproducts include fuel gas, naphtha and low-pour diesel. ��
��
��
Description: MSDW is targeted for hydrocracked or severely hydrotreated ��
stocks. The improved selectivity of MSDW for the highly isoparaffinic-lube ��
��
components results in higher lube yields and VIs. The process uses mul- ��
tiple catalyst systems with multiple reactors. Internals are proprietary (the ��
��
Spider Vortex Quench Zone technology is used). Feed and recycle gases
are preheated and contact the catalyst in a down-flow-fixed-bed reactor. ��
Reactor effluent is cooled, and the remaining aromatics are saturated in a ��
post-treat reactor. The process can be integrated into a lube hydrocracker ��
or lube hydrotreater. Post-fractionation is targeted for client needs. ��

Operating conditions:
Temperatures, ° F 550 – 800
Hydrogen partial pressures, psig 500 – 2,500
LHSV 0.4 – 3.0
Conversion depends on feed wax content Installation: Eight units are operating and four are in design.
Pour point reduction as needed. Licensor: ExxonMobil Research and Engineering Co.
Yields:
Light neutral Heavy neutral
Lube yield, wt% 94.5 96.5
C1– C4, wt% 1.5 1.0
C5– 400°F, wt% 2.7 1.8
400°F – Lube, wt% 1.5 1.0
H2 cons, scf / bbl 100 – 300 100 – 300

Economics: $3,000 – 5,500 per bpsd installed cost (US Gulf Coast).
Catalytic reforming
Application: Increase the octane of straight-run or cracked naphthas for
��
gasoline production.
��
Products: High-octane gasoline and hydrogen-rich gas. Byproducts may
be LPG, fuel gas and steam.
��
Description: Semi-regenerative multibed reforming over platinum or bi-
metallic catalysts. Hydrogen recycled to reactors at the rate of 3 mols / ��
mol to 7 mols /mol of feed. Straight-run and /or cracked feeds are typi- ��
cally hydrotreated, but low-sulfur feeds (<10 ppm) may be reformed
without hydrotreatment. ��
��
��
Operating conditions: 875°F to 1,000°F and 150 psig to 400 psig reac- ��
��
tor conditions. ��
��
Yields: Depend on feed characteristics, product octane and reactor pres-
sure. The following yields are one example. The feed contains 51.4%
parafﬁns, 41.5% naphthenes and 7.1% aromatics, and boils from 208°F
to 375°F (ASTM D86). Product octane is 99.7 RONC and average reactor
pressure is 200 psig.
Economics:
Utilities, (per bbl feed)
Fuel, 103 Btu release 275
Component wt% vol%
Electricity, kWh 7.2
H2 2.3 1,150 scf/bbl
Water, cooling (20°F rise), gal 216
C1 1.1 —
Steam produced (175 psig sat), lb 100
C2 1.8 —
C3 3.2 — Licensor: CB&I Howe-Baker.
iC4 1.6 —
nC4 2.3 —
C5+ 87.1 —
LPG — 3.7
Reformate — 83.2
Catalytic reforming
Application: The CCR Platforming process is used throughout the world ��
in the petroleum and petrochemical industries. It produces feed for an ��
��
aromatics complex or a high-octane gasoline blending product and a ��
��
signiﬁcant amount of hydrogen. ��
��
��
Description: Hydrotreated naphtha feed is combined with recycle hy- ��

drogen gas and heat exchanged against reactor efﬂuent. The combined
feed is then raised to reaction temperature in the charge heater and sent ��

to the reactor section. ��

Radial-flow reactors are arranged in a vertical stack. The predomi-
��
nant reactions are endothermic; so an interheater is used between each ��
��
reactor to reheat the charge to reaction temperature. The effluent from
��
the last reactor is heat exchanged against combined feed, cooled and ��
��
��
split into vapor and liquid products in a separator. The vapor phase is
hydrogen-rich. A portion of the gas is compressed and recycled back to
the reactors. The net hydrogen-rich gas is compressed and charged to-
gether with the separator liquid phase to the product recovery section.
This section is engineered to provide optimum performance.
Catalyst flows vertically by gravity down the reactor stack. Over
time, coke builds up on the catalyst at reaction conditions. Partially de-
activated catalyst is continually withdrawn from the bottom of the reac-
tor stack and transferred to the CCR regenerator.

Installation: UOP commercialized the CCR Platforming process in 1971

and now has commissioned more than 180 units (more than 3.9 million
bpd of capacity) with another 30 in various stages of design, construc-
tion and commissioning.

Efficiency/product quality: Commercial onstream efﬁciencies of more

than 95% are routinely achieved in CCR Platforming units.

Licensor: UOP LLC.

Coking
Application: Conversion of atmospheric and vacuum residuals, hydro- ��
treated and hydrocracked resids, visbroken resids, asphalt, pyrolysis tar,
��
decant oil, coal tar, pitch, solvent-reﬁned and Athabasca bitumen. �
� � ��
Description: Feedstock is introduced (after heat exchange) to the bot-
tom of the coker fractionator (1) where it mixes with condensed recycle.
The mixture is pumped through the coker heater (2) where the desired ��
coking temperature is achieved, to one of two coke drums (3). Steam ��
or boiler feedwater is injected into the heater tubes to prevent coking in ��
the furnace tubes. Coke drum overhead vapors ﬂow to the fractionator �
��
(1) where they are separated into an overhead stream containing the
��
wet gas, LPG and naphtha; two gas oil sidestreams; and the recycle that
rejoins the feed. ��
The overhead stream is sent to a vapor recovery unit (4) where the ��
��
individual product streams are separated. The coke that forms in one of
at least two (parallel connected) drums is then removed using high-pres-
sure water. The plant also includes a blow-down system, coke handling
and a water recovery system.

Operating conditions: Products, wt%

Heater outlet temperature, °F 900 –950 Gas + LPG 7.9 9.0 3.9
Coke drum pressure, psig 15 –90 Naphtha 12.6 11.1 —
Recycle ratio, vol/vol feed, % 0 –100 Gas oils 50.8 44.0 31.0
Coke 28.7 35.9 65.1
Yields:
Vacuum residue of Economics:
Middle East hydrotreated Coal tar Investment (basis: 20,000 bpsd straight-run vacuum residue feed,
Feedstock vac. residue bottoms pitch US Gulf Coast 2006, fuel-grade coke, includes vapor recovery), US$
Gravity, °API 7.4 1.3 –21.0 per bpsd (typical) 6,000
Sulfur, wt% 4.2 2.3 0.5

Conradson
carbon, wt% 20.0 27.6 48
Continued 
Coking, continued
Economics (continued):
Utilities, typical/bbl of feed:
Fuel, 103 Btu 123
Electricity, kWh 3.6
Steam (exported), lb 1
Water, cooling, gal 58
Boiler feedwater, lbs 38
Condensate (exported), lbs 24

Installation: More than 60 units.

Reference: Mallik, R. and G. Hamilton, “Delayed coker design consid-
erations and project execution,” NPRA 2002 Annual Meeting, March
17–19, 2002.

Licensor: ABB Lummus Global.

Coking
Application: Upgrading of petroleum residues (vacuum residue, bitumen,
solvent-deasphalter pitch and fuel oil) to more valuable liquid products ��
(LPG, naphtha, distillate and gas oil). Fuel gas and petroleum coke are ��
also produced.
��
Description: The delayed coking process is a thermal process and con- ��
��
sists of fired heater(s), coke drums and main fractionator. The cracking ��
and coking reactions are initiated in the fired heater under controlled
time-temperature-pressure conditions. The reactions continue as the
��
process stream moves to the coke drums. Being highly endothermic,
the coking-reaction rate drops dramatically as coke-drum temperature ��
��
decreases. Coke is deposited in the coke drums. The vapor is routed to
the fractionator, where it is condensed and fractionated into product
streams—typically fuel gas, LPG, naphtha, distillate and gas oil.
When one of the pair of coke drums is full of coke, the heater outlet
stream is directed to the other coke drum. The full drum is taken offline,
cooled with steam and water and opened. The coke is removed by hy-
draulic cutting. The empty drum is then closed, warmed-up and made
ready to receive feed while the other drum becomes full.
ConocoPhillips ThruPlus Delayed Coking Technology provides sev- • Optimum heater design and operation. The proprietary design
eral advantages: methodology minimizes the conditions that cause coke deposits in the
• Experienced owner-operator. More coke has been processed heater tubes. When combined with ConocoPhillips’ patented distillate
by ConocoPhillips’ ThruPlus Delayed Coking Technology than by any recycle technology, the result is maximum furnace run-length and im-
other competing process. The company has more than 50+ years ex- proved liquid yields.
perience and today owns and operates 17 delayed cokers worldwide, • Highly reliable coke handling system. The sloped concrete
with a combined capacity of more than 650,000 bpd. First licensing our wall and pit-pad system allows space for an entire drum of coke to be
ThruPlus Delayed Coking Technology in 1981, there are now 31 installa- cut without moving any of it. This allows the process side to operate at
tions worldwide, with a combined capacity of 1.1 million bpd and more full capacity, even if there is an issue with the coke handling system. The
coming online in the US, Canada, and Brazil. sloped wall also improves safety and requires little maintenance over the
• Robust coke drum design. Drums designed to ConocoPhillips’ life of the unit.
specifications provide long operating service life—more than 20 years • Short coke drum cycle time. We push the limits of reducing
without severe bulging or cracking—even when run on short drum cy- cycle times, running sustained 10-hour cycles, producing both anode
cles. The key is understanding drum fatigue at elevated temperatures. and fuel coke. Success in safely reducing cycle time requires a thorough
We do understand, and we design accordingly. understanding of each phase of the cycle and the process. Since we’ve
Coking, continued Economics: The economic benefits of the increase in liquid yields of-
fered by the ConocoPhillips process are substantial. With a very conser-
been cycling drums since 1953, we understand every step. vative estimated price of $1.50 US/gallon of transportation fuels, annual
• Shot coke handling. The design and operating procedures allow earnings for a 2% increase in liquid yields can exceed US$15 million for
minimizing risks associated with handling shot coke, while maximizing a 40,000-bpd day coker.
profits.
Installation: Low investment cost and attractive yield structure has made
Other distinguishing features that improve emissions include:
delayed coking the technology of choice for bottom-of-the-barrel up-
• Advanced closed blowdown system, virtually eliminates hy-
grading. Numerous delayed coking units are operating in petroleum re-
drocarbon discharge, condenses most steam and recovers water. The
fineries worldwide.
recovered water is returned to the drilling and quench system.
• Dust suppression is achieved by surrounding the coke storage Licensor: ConocoPhillips.
area with high walls and eliminates wheeled equipment. The coke is
handled by overhead crane—a safer alternative.
• Reuse of drilling and quench water uses an effective fines re-
moval system to improve water quality.
• Processing of oil bearing solid waste. With a widely-used tech-
nology, oily solids are recovered. The oil is contained within the coker
process, and the solids are combined with the coke.
Coking ��

� ��
Application: Upgrade residues to lighter hydrocarbon fractions using the
Selective Yield Delayed Coking (SYDEC) process.
� �
Description: Charge is fed directly to the fractionator (1) where it com- ��
bines with recycle and is pumped to the coker heater. The mixture is
heated to coking temperature, causing partial vaporization and mild
cracking. The vapor-liquid mix enters a coke drum (2 or 3) for further � ��
cracking. Drum overhead enters the fractionator (1) to be separated into ��
gas, naphtha, and light and heavy gas oils. Gas and naphtha enter the
vapor recovery unit (VRU)(4). There are at least two coking drums, one
��
coking while the other is decoked using high-pressure water jets. The ��
coking unit also includes a coke handling, coke cutting, water recovery
��
and blowdown system. Vent gas from the blowdown system is recov-
ered in the VRU.

Operating conditions: Typical ranges are: Utilities, typical per bbl feed:
Heater outlet temperature, ºF 900 – 950 Fuel, 103 Btu 120
Coke drum pressure, psig 15 – 100 Electricity, kWh 3
Recycle ratio, equiv. fresh feed 0 – 1.0 Steam (exported), lb 35
Increased coking temperature decreases coke production; increases Water, cooling, gal 36
liquid yield and gas oil end point. Increasing pressure and/or recycle ra-
Installations: Currently, 52 delayed cokers are installed worldwide with
tio increases gas and coke make, decreases liquid yield and gas oil end
a total installed capacity over 2.5 million bpsd
point.
References: Handbook of Petroleum Refining Processes, Third Ed., Mc-
Yields:
Graw-Hill, pp. 12.33 –12.89.
Operation:
“Delayed coking revamps,” Hydrocarbon Processing, September 2004.
Products, wt% Max dist. Anode coke Needle coke
“Residue upgrading with SYDEC Delayed Coking: Benefits & Eco-
Gas 8.7 8.4 9.8
nomics,” AIChE Spring National Meeting, April 23–27, 2006, Orlando.
Naphtha 14.0 21.6 8.4
“Upgrade refinery residuals into value-added products,” Hydrocar-
Gas oil 48.3 43.8 41.6
bon Processing, June 2006.
Coke 29.3 26.2 40.2
Licensor: Foster Wheeler/UOP LLC.
Economics:
Investment (basis 65,000 –10,000 bpsd)
2Q 2005 US Gulf), $ per bpsd 3,000 –5,200
Coking, fluid
Application: Continuous fluid, bed coking technology to convert heavy
hydrocarbons (vacuum residuum, extra heavy oil or bitumen) to full- Reactor products
range lighter liquid products and fluid coke. Product coke can be sold to fractionator
as fuel or burned in an integrated fluid bed boiler to produce steam and Flue gas to CO boiler
power.
START 1
Products: Liquid products can be upgraded through conventional 3
hydrotreating. Fluid coke is widely used as a solid fuel, with particular 2
Net coke
advantages in cement kilns and in fluid-bed boilers.
Air
Description: Feed (typically 1,050°F+ vacuum resid) enters the scrubber blower

(1) for heat exchange with reactor overhead effluent vapors. The scrub- Air
ber typically cuts 975°F+ higher boiling reactor effluent hydrocarbons Cold Hot
coke coke
for recycle back to the reactor with fresh feed. Alternative scrubber con-
figurations provide process flexibility by integrating the recycle stream
with the VPS or by operating once-through which produces higher liq-
uid yields. Lighter overhead vapors from the scrubber are sent to con-
ventional fractionation and light ends recovery. In the reactor (2), feed
is thermally cracked to a full range of lighter products and coke.
The heat for the thermal cracking reactions is supplied by circulating C5+ liquids, wt% 58.1 62.3
coke between the burner (3) and reactor (2). About 20% of the coke
is burned with air to supply process heat requirements, eliminating the Net product coke, wt % 25.7 23.9
need for an external fuel supply. The rest of the coke is withdrawn and Coke consumed for heat, wt% 4.4 3.4
either sold as a product or burned in a fluid bed boiler. Properties of the
fluid coke enable ease of transport and direct use in fuel applications, Investment: TPC, US Gulf Coast, 2Q 2003 estimate including gas pro-
including stand alone or integrated cogeneration facilities. cessing, coke handling and wet gas scrubbing for removing SOx from
the burner overhead
Yields: Example, typical Middle East vacuum resid (~25 wt% Concar- Capital investment, $/bp/sd 3,300
bon, ~5 wt% sulfur):
Recycle Once-Through Competitive advantages:
Light ends, wt% 11.8 10.4 • Single train capacities >100 Mbpsd; greater than other processes
Naphtha (C5-350°F), wt% 11.5 9.5 • Process wide range of feeds, especially high metals, sulfur and CCR
Distillate (350 – 650°F), wt% 14.5 13.1 • Internally heat integrated, minimal use of fuel gas, and lower coke
Heavy gas oil (650°F+), wt% 32.1 39.7 production than delayed coking
Continued 
Coking, fluid, continued
• Lower investment and better economy of scale than delayed cok-
ing
• Efficient integration with fluid bed boilers for cogeneration of
steam and electric power.

Licensor: ExxonMobil Research and Engineering Co.

Coking, flexi
Application: Continuous fluid-bed coking technology to convert heavy Low heating
Tertiary cyclones value coke gas
hydrocarbons (vacuum residuum, extra heavy oil or bitumen) to full-
range lighter liquid products and Flexigas—a valuable lower Btu fuel Reactor products Steam
to fractionator generation
gas. Applicable for complete conversion of resid in refineries with lim- 5
ited outlets for coke, for heavy feed conversion at the resource, and 5 Direct
locations where low-cost clean fuel is needed or where natural gas has contact
START 1 cooler
high value. Sour
water
2
Products: Liquid products can be upgraded through conventional 3 4
Coke
fines
hydrotreating. Clean Flexigas with <10 vppm H2S can be burned in fur-
naces or boilers, replacing fuel oil, fuel gas or natural gasfuels.
Steam
Description: FLEXICOKING has essentially the same reactor (1)/scrubber Hot Air
Cold blower
(2) sections as FLUID COKING, and also has the same process flexibility coke coke
options: recycle, once-through and VPS integrated. Air
Process heat for the coking and gasification reactions is supplied
by circulating hot coke between the heater (3) the reactor (2), and the
gasifier (4). Coke reacts with air and steam in the gasifier (4) to produce
heat and lower BTU gas that is cleaned with FLEXSORB hinder amine
treating to <10 vppm of H2S. About 97% of the coke generated is con- Purge coke, wt% ~1 ~1
sumed in the process; a small amount of purge coke is withdrawn from Flexigas, MBtu/kbbl of feed 1,200 1,100
the heater (3) and fines system (5), which can be burned in cement kilns
or used for vanadium recovery. Partial gasification/coke withdrawal and Investment: TPC, US Gulf Coast, 2Q03 estimate including gas process-
oxygen-enrichment can be used to provide additional process flexibility. ing, coke handling, Flexigas treating and distribution
Capital investment, $/bp/sd 4,700
Yields: Example, typical Middle East vacuum resid (~25 wt% Concar-
bon, ~5 wt% sulfur): Competitive advantages:
Recycle Once-through • Fully continuous commercially proven integrated fluid bed coking
Light ends, (C4 ),wt% 11.8 10.4 and fluid bed gasification process that produces valuable liquid products
Naphtha (C5-350°F), wt% 11.5 9.5 and gaseous fuels.
Distillate (350 – 650°F), wt% 14.5 13.1 • Low value coke is converted in the process to clean product Flexi-
Heavy gas oil (650°F+), wt% 32.1 39.7 gas fuel gas for use within the refinery or by nearby third-party power
C5+ Liquids, wt% 58.1 62.3 plants or other consumers.
Continued 
Coking, flexi, continued
• Particularly attractive for SAGD tar sands upgrading with large fuel
requirements. Much lower investment and more reliable than delayed
coking plus partial oxidation or direct gasification of solids or heavy
feeds.

Licensor: ExxonMobil Research and Engineering Co.

Crude distillation
Application: Separates and recovers the relatively lighter fractions (e.g., 6
Flash gas
naphtha, kerosine, diesel and cracking stock) from a fresh crude oil charge. Light naphtha
The vacuum flasher processes the crude distillation bottoms to produce Heavy naphtha
an increased yield of liquid distillates and a heavy residual material. 7 Kerosine
5 8
Description: The charge is preheated (1), desalted (2) and directed to a 4 Diesel
3 9
preheat train (3) where it recovers heat from product and reflux streams. Cracker feed
The typical crude fired heater (4) inlet temperature is on the order of 10
To vac. system
550 ° F, while the outlet temperature is on the order of 675°F to 725°F. 2
Lt. vac. gas oil
Heater effluent then enters a crude distillation column (5) where light
Stm. Stm. Hvy. vac.
naphtha is drawn off the tower overhead (6); heavy naphtha, kerosine, 1 12 gas oil
diesel and cracking stock are sidestream drawoffs. External reflux for Crude 11
Vac. gas oil
the tower is provided by pumparound streams (7–10). The atmospheric Stm. (cracker feed)
START
residue is charged to a fired heater (11) where the typical outlet tem-
Asphalt
perature is on the order of 725 °F to 775°F.
From the heater outlet, the stream is fed into a vacuum tower
(12), where the distillate is condensed in two sections and withdrawn
as two sidestreams. The two sidestreams are combined to form crack-
ing feedstock. An asphalt base stock is pumped from the bottom of Economics:
the tower. Two circulating reflux streams serve as heat removal media Investment ( basis: 100,000–50,000 bpsd,
for the tower. 2nd Q, 2005, US Gulf ), $ per bpsd 900 – 1,400
Yields: Typical for Merey crude oil: Utility requirements, typical per bbl fresh feed
Steam, lb 24
Crude unit products wt% °API Pour, °F
Fuel (liberated), 10 3 Btu ( 80 – 120 )
Overhead & naphtha 6.2 58.0 —
Power, kWh 0.6
Kerosine 4.5 41.4 – 85
Water, cooling, gal 300 – 400
Diesel 18.0 30.0 – 10
Gas oil 3.9 24.0 20 Installation: Foster Wheeler has designed and constructed crude units
Lt. vac. gas oil 2.6 23.4 35 having a total crude capacity in excess of 15 MMbpsd.
Hvy. vac. gas oil 10.9 19.5 85 Reference: Encyclopedia of Chemical Processing and Design, Marcel-
Vac. bottoms 53.9 5.8 (120)* Dekker, 1997, pp. 230 – 249.
Total 100.0 8.7 85
Licensor: Foster Wheeler.
*Softening point, °F
Note: Crude unit feed is 2.19 wt% sulfur. Vacuum unit feed is 2.91 wt% sulfur.
Crude distillation
��
Application: The Shell Bulk CDU is a highly integrated concept. It sepa-
rates the crude in long residue, waxy distillate, middle distillates and a
naphtha minus fraction. Compared with stand-alone units, the overall ��
��
integration of a crude distillation unit (CDU), hydrodesulfurization unit ��
(HDS), high vacuum unit (HVU) and a visbreaker (VBU) results in a 50% �
� ��
reduction in equipment count and significantly reduced operating costs. � ��
��
A prominent feature embedded in this design is the Shell deepflash HVU � ��
technology. This technology can also be provided in cost-effective pro- ��
cess designs for both feedprep and lube oil HVUs as stand-alone units. ��
� ��
For each application, tailor-made designs can be produced. ��
��
�
Description: The basic concept of the bulk CDU is the separation of ��
��
the naphtha minus and the long residue from the middle distillate ��
fraction, which is routed to the HDS. After desulfurization in the HDS ��
unit, final product separation of the bulk middle distillate stream from
the CDU takes place in the HDS fractionator (HDF), which consists of a
main atmospheric fractionator with side strippers.
The long residue is routed hot to a feedprep HVU, which recovers
the waxy distillate fraction from long residue as the feedstock for a Economics: Due to the incorporation of Shell high capacity internals
cat-cracker or hydrocracker unit (HCU). Typical flashzone conditions and the deeply integrated designs, an attractive CAPEX reduction can be
are 415°C and 24 mbara. The Shell design features a deentrainment achieved. Investment costs are dependent on the required configuration
section, spray sections to obtain a lower flashzone pressure, and a and process objectives.
VGO recovery section to recover up to 10 wt% as automotive diesel. Installation: Over 100 Shell CDUs have been designed and operated
The Shell furnace design prevents excessive cracking and enables a since the early 1900s. Additionally, a total of some 50 HVU units have
5-year run length between decoke. been built while a similar number has been debottlenecked, including
Yields: Typical for Arabian light crude many third-party designs of feedprep and lube oil HVUs.
Products wt, % Licensor: Shell Global Solutions International B.V.
Gas C1 – C4 0.7
Gasoline C5 – 150°C 15.2
Kerosine 150 – 250°C 17.4
Gasoil (GO) 250 – 350°C 18.3
VGO 350 – 370°C 3.6
Waxy distillate (WD) 370 – 575°C 28.8
Residue 575°C+ 16.0
Crude distillation
Application: The D2000 process is progressive distillation to minimize the ��
��
total energy consumption required to separate crude oils or condensates ��
��
into hydrocarbon cuts, which number and properties are optimized to fit
� ��
with sophisticated refining schemes and future regulations. This process is
applied normally for new topping units or new integrated topping/vacu- ��
um units but the concept can be used for debottlenecking purpose. ��
��
�
Products: This process is particularly suitable when more than two ��
� �
naphtha cuts are to be produced. Typically the process is optimized to �
��
produce three naphtha cuts or more, one or two kerosine cuts, two at-
mospheric gas oil cuts, one vacuum gas oil cut, two vacuum distillates � ��
cuts, and one vacuum residue. ��

Description: The crude is preheated and desalted (1). It is fed to a first ��
��
dry reboiled pre-flash tower (2) and then to a wet pre-flash tower (3).
The overhead products of the two pre-flash towers are then fraction-
ated as required in a gas plant and rectification towers (4).
The topped crude typically reduced by 2/3 of the total naphtha cut is
then heated in a conventional heater and conventional topping column
(5). If necessary the reduced crude is fractionated in one deep vacuum Utility requirements, typical per bbl of crude feed:
column designed for a sharp fractionation between vacuum gas oil, two Fuel fired, 103 btu 50–65
vacuum distillates (6) and a vacuum residue, which could be also a road Power, kWh 0.9–1.2
bitumen. Steam 65 psig, lb 0–5
Extensive use of pinch technology minimizes heat supplied by heat- Water cooling, (15°C rise) gal 50–100
ers and heat removed by air and water coolers. Total primary energy consumption:
This process is particularly suitable for large crude capacity from for Arabian Light or Russian Export Blend: 1.25 tons of fuel
150,000 to 250,000 bpsd. per 100 tons of Crude
It is also available for condensates and light crudes progressive distil- for Arabian Heavy 1.15 tons of fuel
lation with a slightly adapted scheme. per 100 tons of Crude
Economics: Installation: Technip has designed and constructed one crude unit and
Investment (basis 230,000 bpsd including atmospheric and one condensate unit with the D2000 concept. The latest revamp proj-
vacuum distillation, gas plant and rectification tower) $750 to
$950 per bpsd (US Gulf Coast 2000).
Continued 
Crude distillation, continued
ect currently in operation shows an increase of capacity of the existing
crude unit of 30% without heater addition.

Licensor: TOTAL and Technip.

Crude topping units
Application: Crude topping units are typically installed in remote areas ��
to provide fuel for local consumption, often for use at pipeline pumping
stations and at production facilities. ��

Products: Diesel is typically the desired product, but kerosine, turbine ��
fuel and naphtha are also produced. ��

Description: Crude topping units comprise of four main sections: pre-

heat/heat recovery section, ﬁred heater, crude fractionation (distillation), ��
��
and product cooling and accumulation. The ﬁred heater provides heat
��
for the plant. Fuel for the heater can be residual products, offgas, natu-
ral gas, distillate product, or combinations of these fuels, depending
on the installation. Heat integration reduces emissions and minimizes ��
process-energy requirements. Depending on the individual site, an elec- ��
trostatic desalter may be required to prevent fouling and plugging, and
control corrosion in the fractionation section.
Crude topping units are modularized, which reduces construction
cost and complexity. Modular units also allow installation in remote ar-
eas with minimal mobilization. These units are typically custom designed
to meet individual customer requirements.
Crude topping units are self-contained, requiring few utilities for opera-
tion. Utility packages, wastewater treatment facilities and other associated
offsites are often supplied, depending on the individual site requirements.

Operating conditions:
Column pressure, psig 0 – 20
Temperature, °F 550 – 650

Licensor: CB&I Howe-Baker.

Deasphalting
Application: Prepare quality feed for FCC units and hydrocrackers from
vacuum residue, and blending stocks for lube oil and asphalt manufac-
turing.

Products: Deasphalted oil (DAO) for catalytic cracking and hydrocracking ��
�
��
feedstocks, resins for specification asphalts, and pitch for specification ��
��
asphalts and residue fuels.
��
Description: Feed and light paraffinic solvent are mixed and then ��
charged to the extractor (1). The DAO and pitch phases, both containing
solvents, exit the extractor. The DAO and solvent mixture is separated � � ��
under supercritical conditions (2). Both the pitch and DAO products are
��
stripped of entrained solvent (3,4). A second extraction stage is utililized
if resins are to be produced.
��

Operating conditions: Typical ranges are:

Solvent various blends of C3– C7 hydrocarbons includ-
ing light naphthas
Pressure, psig 300 – 600
Temp., °F 120 – 450 Ni/V, wppm 0.25/0.37 4.5/10.3
Solvent to oil ratio: 4/1 to 13/1 Pitch
Softening point, R&B, ºF 149 240
Yields:
Penetration@77ºF 12 0
Feed, type Lube oil Cracking stock
Gravity, ºAPI 6.6 6.5 Economics:
Sulfur, wt% 4.9 3.0 Investment (basis: 40,000 –2,000 bpsd)
CCR, wt% 20.1 21.8 2Q 2005, US Gulf, $/bpsd 800 – 3,000
Visc, SSU@210ºF 7,300 8,720 Utilities, typical per bbl feed:
Ni/V, wppm 29/100 46/125 Fuel, 103 Btu (hot oil) 56 – 100
DAO Electricity, kWh 1.9 – 2.0
Yield, vol.% of feed 30 65 Steam, 150 psig, lb 6–9
Gravity, ºAPI 20.3 15.1 Water, cooling (25ºF rise), gal 10
Sulfur, wt% 2.7 2.2
CCR, wt% 1.4 6.2
Visc., SSU@210ºF 165 540 Continued 
Deasphalting, continued
Installations: Over 50 units installed; this also includes both UOP and
Foster Wheeler units originally licensed separately before merging the
technologies in 1996.

References: Handbook of Petroleum Reﬁning Processes, Third Ed., Mc-

Graw Hill, 2003, pp. 10.37–10.61.
“When Solvent Deasphalting is the Most Appropriate Technology
for Upgrading Residue,” International Downstream Technology Confer-
ence, February 15 –16, 2006, London.

Licensor: Foster Wheeler/UOP LLC.

Deep catalytic cracking
Application: Selective conversion of gasoil and parafﬁnic residual feed-
stocks.

Products: C2– C5 oleﬁns, aromatic-rich, high-octane gasoline and

distillate.

Description: DCC is a ﬂuidized process for selectively cracking a wide

variety of feedstocks to light olefins. Propylene yields over 24 wt% are
achievable with paraffinic feeds. A traditional reactor/regenerator unit
design uses a catalyst with physical properties similar to traditional FCC
catalyst. The DCC unit may be operated in two operational modes: max-
imum propylene (Type I) or maximum iso-olefins (Type II).
Each operational mode utilizes unique catalyst as well as reaction
conditions. Maximum propylene DCC uses both riser and bed cracking
at severe reactor conditions while Type II DDC uses only riser cracking
like a modern FCC unit at milder conditions.
The overall flow scheme of DCC is very similar to that of a conven-
tional FCC. However, innovations in the areas of catalyst development,
process variable selection and severity and gas plant design enables Products (wt% of fresh feed) DCC Type I DCC Type II FCC
Ethylene 6.1 2.3 0.9
the DCC to produce significantly more olefins than FCC in a maximum
Propylene 20.5 14.3 6.8
olefins mode of operation.
Butylene 14.3 14.6 11.0
This technology is quite suitable for revamps as well as grassroot
in which IC4= 5.4 6.1 3.3
applications. Feed enters the unit through proprietary feed nozzles, as
Amylene — 9.8 8.5
shown in the schematic. Integrating DCC technology into existing refin-
in which IC5= — 6.5 4.3
eries as either a grassroots or revamp application can offer an attractive
opportunity to produce large quantities of light olefins. Installation: Six units are currently operating in China and one in Thailand.
In a market requiring both propylene and ethylene, use of both Several more units are under design in China and one in Saudi Arabia.
thermal and catalytic processes is essential, due to the fundamental
differences in the reaction mechanisms involved. The combination of Reference: Dharia, D., et al., “Increase light olefins production,” Hydro-
thermal and catalytic cracking mechanisms is the only way to increase carbon Processing, April 2004, pp. 61–66.
total olefins from heavier feeds while meeting the need for an increased Licensor: Shaw Stone & Webster and Research Institute of Petroleum
propylene to ethylene ratio. The integrated DCC/steam cracking com- Processing, Sinopec.
plex offers significant capital savings over a conventional stand-alone
refinery for propylene production.
Deep thermal conversion
��
Application: The Shell Deep Thermal Conversion process closes the gap
between visbreaking and coking. The process yields a maximum of distil- ��
lates by applying deep thermal conversion of the vacuum residue feed
and by vacuum flashing the cracked residue. High-distillate yields are
obtained, while still producing a stable liquid residual product, referred �
to as liquid coke. The liquid coke, not suitable for blending to commer- ��
cial fuel, is used for speciality products, gasification and/or combustion,
e.g., to generate power and/or hydrogen. ��
�
Description: The preheated short residue is charged to the heater (1) ��
�
and from there to the soaker (2), where the deep conversion takes place. �
The conversion is maximized by controlling the operating temperature ��
� ��
and pressure. The soaker effluent is routed to a cyclone (3). The cyclone ��
overheads are charged to an atmospheric fractionator (4) to produce the
desired products like gas, LPG, naphtha, kero and gasoil. The cyclone
and fractionator bottoms are subsequently routed to a vacuum flasher
(5), which recovers additional gasoil and waxy distillate. The residual liq-
Utilities, typical consumption/production for a 25,000-bpd unit,
uid coke is routed for further processing depending on the outlet.
dependent on configuration and a site’s marginal econmic values for
Yields: Depend on feed type and product specifications. steam and fuel:
Fuel as fuel oil equivalent, bpd 417
Feed, vacuum residue Middle East Power, MW 1.2
Viscosity, cSt @100°C 615 Net steam production (18 bar), tpd 370
Products in % wt. on feed
Gas 3.8 Installation: To date, six Shell Deep Thermal Conversion units have been
Gasoline, ECP 165°C 8.2 licensed. In four cases, this has involved revamping an existing Shell
Gas oil, ECP 350°C 19 Soaker Visbreaker unit. Post startup services and technical services for
Waxy distillate, ECP 520°C 22.8 existing units are available from Shell Global Solutions.
Residue ECP 520°C+ 46.2
Reference: Hydrocarbon Engineering, September 2003.
Economics: The typical investment for a 25,000-bpd unit will be about
$1,900 to $2,300/bbl installed, excluding treating facilities. (Basis: West- Licensor: Shell Global Solutions International B.V. and ABB Lummus
ern Europe, 2004.). Global B.V.
Desulfurization
Application: GT-BTXPlus addresses overall plant profitability by desul-
furizing the FCC stream with no octane loss and decreased hydrogen ��
consumption by using a proprietary solvent in an extractive distillation ��
��
system. This process also recovers valuable aromatics compounds. Olefin-
rich raffinate can be recycled to FCC or to aromatizing unit. ��
��
� ��
��
Description: FCC gasoline, with endpoint up to 210 °C, is fed to the GT- ��
��
BTXPlus unit, which extracts sulfur and aromatics from the hydrocarbon �
stream. The sulfur and aromatic components are processed in a conven-
tional hydrotreater to convert the sulfur into H 2S. Because the portion ��
of gasoline being hydrotreated is reduced in volume and free of olefins, ��

hydrogen consumption and operating costs are greatly reduced. In con-

trast, conventional desulfurization schemes process the majority of the
gasoline through hydrotreating and caustic-washing units to eliminate ��

the sulfur. That method inevitably results in oleﬁn saturation, octane

downgrade and yield loss.
GT-BTXPlus has these advantages:
• Segregates and eliminates FCC-gasoline sulfur species to meet a tionally caustic washed before routing to the gasoline pool, or to aro-
pool gasoline target of 20 ppm matizing unit.
• Preserves more than 90% of the olefins from being hydrotreated Rich solvent, containing aromatics and sulfur compounds, is routed
in the HDS unit; and thus, prevents significant octane loss and to the solvent recovery column, (SRC), where the hydrocarbons and sul-
reduces hydrogen consumption fur species are separated, and lean solvent is recovered in columns bot-
• Fewer components are sent to the HDS unit; consequently, a toms. The SRC overhead is hydrotreated by conventional means and
smaller HDS unit is needed and there is less yield loss used as desulfurized gasoline, or directed to an aromatics production
• High-purity BTX products can be produced from the aromatic-rich plant. Lean solvent from the SRC bottoms are treated and recycled back
extract stream after hydrotreating to the EDC.
• Olefin-rich raffinate stream (from the ED unit) can be recycled to
the FCC unit to increase the light olefin production. Economics: Estimated installed cost of $1,000 / bpd of feed and produc-
FCC gasoline is fed to the extractive distillation column (EDC). In tion cost of $0.50 / bbl of feed for desulfurization and dearomatization.
a vapor-liquid operation, the solvent extracts the sulfur compounds
Licensor: GTC Technology Inc.
into the bottoms of the column along with the aromatic components,
while rejecting the olefins and nonaromatics into the overhead as raf-
finate. Nearly all of the nonaromatics, including olefins, are effectively
separated into the raffinate stream. The raffinate stream can be op-
Dewaxing
Application: Bechtel’s Dewaxing process is used to remove waxy com-
ponents from lubrication base-oil streams to simultaneously meet de- � �
sired low-temperature properties for dewaxed oils and produce slack ��
wax as a byproduct.

Description: Waxy feedstock (rafﬁnate, distillate or deasphalted oil) is ��

mixed with a binary-solvent system and chilled in a very closely controlled �
manner in scraped-surface double-pipe exchangers (1) and refrigerated
chillers (2) to form a wax/oil/solvent slurry.
The slurry is filtered through the primary filter stage (3) and dewaxed
oil mixture is routed to the dewaxed oil recovery section (5) to separate ��
� �
� ��
solvent from oil. Prior to solvent recovery, the primary filtrate is used to ��
cool the feed/solvent mixture (1). Wax from the primary stage is slurried ��
with cold solvent and filtered again in the repulp filter (4) to reduce the ��
��
oil content to approximately 10%.
The repulp filtrate is reused as dilution solvent in the feed chilling
train. The wax mixture is routed to a solvent-recovery section (6) to re-
move solvent from the product streams (hard wax and soft wax). The
recovered solvent is collected, dried (7) and recycled back to the chilling
and filtration sections.

Economics:
Investment (Basis: 7,000-bpsd feedrate
capacity, 2006 US Gulf Coast), $/bpsd 11,200
Utilities, typical per bbl feed:
Fuel, 103 Btu (absorbed) 160
Electricity, kWh 15
Steam, lb 35
Water, cooling (25°F rise), gal 1,100

Installation: Over 100 have been licensed and built.

Licensor: Bechtel Corp.
Dewaxing
��
Application: Selectively convert feedstock’s waxy molecules by isomeriza-
tion in the presence of ISODEWAXING Catalysts. The products are high-
quality base oils that can meet stringent cold ﬂow properties. ��

Description: ISODEWAXING Catalysts are very special catalysts that con- �

��
� �
vert feedstocks with waxy molecules (containing long, paraffinic chains) ��
into two or three main branch isomers that have low-pour points. The
product also has low aromatics content. Typical feeds are: raffinates,
slack wax, foots oil, hydrotreated VGO, hydrotreated DAO and uncon- ��
�
verted oil from hydrocracking. � ��
��
As shown in the simplified flow diagram, waxy feedstocks are mixed
with recycle hydrogen and fresh makeup hydrogen, heated and charged ��
to a reactor containing ISODEWAXING Catalyst (1). The effluent will
have a much lower pour point and, depending on the operating severity,
the aromatics content is reduced by 50– 80% in the dewaxing reactor.
In a typical configuration, the effluent from a dewaxing reactor is
cooled down and sent to a finishing reactor (2) where the remaining Utilities: Typical per bbl feed:
single ring and multiple ring aromatics are further saturated by the ISO- Power, kW 3.3
FINISHING Catalysts. The effluent is flashed in high-pressure and low- Fuel , kcal 13.4 x 103
pressure separators (3, 4). Small amounts of light products are recovered Steam, superheated, required, kg 5.3
in a fractionation system (5). Steam, saturated, produced, kg 2.4
Water, cooling, kg 192
Yields: The base oil yields strongly depend on the feedstocks. For a typi-
Chemical-hydrogen consumption, Nm3/m3 oil 30~50
cal low wax content feedstock, the base oil yield can be 90–95%. Higher
wax feed will have a little lower base oil yield. Installation: More than twelve units are in operation and six units are in
various stages of design or construction.
Economics:
Investment: This is a moderate investment process; for a typical size Reference: NPRA Annual Meeting, March 2005, San Francisco, Paper
ISODEWAXING/ISOFINISHING Unit, the capital for ISBL AM-05-39.
is about 6,000 $/bpsd.
Licensor: Chevron Lummus Global LLC.
Dewaxing/wax deoiling
Application: Bechtel’s Dewaxing/Wax Fractionation processes are used
to remove waxy components from lubrication base-oil streams to si- � � �
multaneously meet desired low-temperature properties for dewaxed oils ��
and produce hard wax as a premium byproduct.

Description: Bechtel’s two-stage solvent dewaxing process can be ex- ��

panded to simultaneously produce hard wax by adding a third deoiling � �
stage using the Wax Fractionation process. Waxy feedstock (raffinate,
��
distillate or deasphalted oil) is mixed with a binary-solvent system and
chilled in a very closely controlled manner in scraped-surface double-
pipe exchangers (1) and refrigerated chillers (2) to form a wax/oil/solvent ��
� � ��
�
slurry. ��
��
The slurry is filtered through the primary filter stage (3) and dewaxed
��
oil mixture is routed to the dewaxed oil recovery section (6) to separate ��
��
solvent from oil. Prior to solvent recovery, the primary filtrate is used to ��

cool the feed/solvent mixture (1).

Wax from the primary stage is slurried with cold solvent and filtered
again in the repulp filter (4) to reduce the oil content to approximately
10%. The repulp filtrate is reused as dilution solvent in the feed chilling
train. The low-oil content slack wax is warmed by mixing with warm Utilities, typical per bbl feed:
solvent to melt the low-melting-point waxes (soft wax) and is filtered in Fuel, 103 Btu (absorbed) 230
a third stage filtration (5) to separate the hard wax from the soft wax. Electricity, kWh 25
The hard and soft wax mixtures are each routed to solvent recovery sec- Steam, lb 25
tions (7,8) to remove solvent from the product streams (hard wax and Water, cooling (25°F rise), gal 1,500
soft wax). The recovered solvent is collected, dried (9) and recycled back
to the chilling and filtration sections. Installation: Seven in service.

Economics: Licensor: Bechtel Corp.

Investment (Basis: 7,000-bpsd feedrate
capacity, 2006 US Gulf Coast), $/bpsd 13,200
Diesel—ultra-low-sulfur diesel (ULSD)
Application: Topsøe ULSD process is designed to produce ultra-low-sulfur ��
��
diesel (ULSD) (5–50 wppm S) from cracked and straight-run distillates.
��
By selecting the proper catalyst and operating conditions, the process
can be designed to produce 5 wppm S diesel at low reactor pressures
��
(<500 psig) or at higher reactor pressure when products with improved ��
density, cetane, and polyaromatics are required.
��
Description: Topsøe ULSD process is a hydrotreating process that com-
��
bines Topsøe’s understanding of deep-desulfurization kinetics, high-ac-
tivity catalyst, state-of-the-art reactor internal, and engineering exper- ��
��
tise in the design of new and revamped ULSD units. The ULSD process
can be applied over a very wide range of reactor pressures. ��
��
Our highest activity BRIM catalyst is speciﬁcally formulated with ��
��
high-desulfurization activity and stability at low reactor pressure (~ 500
��
psig) to produce 5 wppm diesel. This catalyst is suitable for revamping
existing low-pressure hydrotreaters or in new units when minimizing
hydrogen consumption.
The highest activity BRIM catalyst is suitable at higher pressure when
secondary objectives such as cetane improvement and density reduction
are required. Topsøe offers a wide range of engineering deliverables to Hoekstra, G., V. Pradhan, K. Knudsen, P. Christensen, I. Vasalos and
meet the needs of the reﬁners. Our offerings include process scoping S. Vousvoukis, “ULSD: Ensuring the unit makes on-spec. product,” NPRA
study, reactor design package, process design package, or engineering Annual Meeting, Salt Lake City, March 2006.
design package.
Licensor: Haldor Topsøe A/S.
Installation: Topsøe has licensed more than 50 ULSD hydrotreaters of
which more than 40 units are designed for less than 10 wppm sulfur in
the diesel. Our reactor internals are installed in more than 60 ULSD units.

References:
Low, G., J. Townsend and T. Shooter, “Systematic approach for the
revamp of a low-pressure hydrotreater to produce 10-ppm, sulfur-free
diesel at BP Conyton Refinery,” 7th ERTC, Paris, November 2002.
Sarup, B., M. Johansen, L. Skyum and B. Cooper, “ULSD Production
in Practice,” 9th ERTC, Prague, November 2004.
Diesel upgrading
Application: Topsøe’s Diesel Upgrading process can be applied for im- ��
provement of a variety of diesel properties, including reduction of diesel ��
specific gravity, reduction of T90 and T95 distillation (Back-end-shift),
reduction of aromatics, and improvements of cetane, cold-flow prop- ��
��
erties, (pour point, clouds point, viscosity and CFPP) and diesel color ��
reduction (poly shift). Feeds can range from blends of straight-run and ��
cracked gas oils up to heavy distillates, including light vacuum gas oil. ��

Description: Topsøe’s Diesel Upgrading process is a combination of treating ��

and upgrading. The technology combines state-of-the-art reactor internals, ��
engineering expertise in quality design, high-activity treating catalyst and ��
��
proprietary diesel upgrading catalyst. Every unit is individually designed to ��
improve the diesel property that requires upgrading. This is done by se-
��
lecting the optimum processing parameters, including unit pressure and ��
LHSV and determining the appropriate Topsøe high-activity catalysts and
plant lay-out. The process is suitable for new units or revamps of existing
hydrotreating units.
In the reactor system, the treating section uses Topsøe’s high-activ-
ity CoMo or NiMo catalyst, such as TK-575 BRIM or TK-576 BRIM, to References: Patel, R., “How are refiners meeting the ultra-low-sulfur
remove feed impurities such as sulfur and nitrogen. These compounds diesel challenge?” NPRA Annual Meeting, San Antonio, March 2003.
limit the downstream upgrading catalyst performance, and the purified Fuente, E., P. Christensen, and M. Johansen, “Options for meeting
stream is treated in the downstream upgrading reactor. Reactor catalyst EU year 2005 fuels specifications,” 4th ERTC, November 1999.
used in the application is dependent on the specific diesel property that
requires upgrading. Reactor section is followed by separation and strip- Installations: A total of 16 units; six in Asia-Pacific region, one in the
ping/fractionation where final products are produced. Middle East, two in Europe and seven HDS/HDA units (see Hydrodearo-
Like the conventional Topsøe hydrotreating process, the diesel up- matization).
grading process uses Topsøe’s graded-bed loading and high-efficiency
patented reactor internals to provide optimal reactor performance and Licensor: Haldor Topsøe A/S.
catalyst utilization. Topsøe’s high-efficiency internals are effective for a
wide range of liquid loading. Topsøe’s graded-bed technology and the
use of shape-optimized inert topping material and catalyst minimize
the pressure drop build-up, thereby reducing catalyst skimming require-
ments and ensuring long catalyst cycle lengths.
Ethers
Application: Production of high-octane reformulated gasoline compo-
nents (MTBE, ETBE, TAME and/or higher molecular-weight ethers) from
��
C1 to C2 alcohols and reactive hydrocarbons in C4 to C6 cuts.

Description: Different arrangements have been demonstrated depend-

ing on the nature of the feeds. All use acid resins in the reaction section.
The process includes alcohol purification (1), hydrocarbon purification ��
� �
(2), followed by the main reaction section. This main reactor (3) operates �
�
under adiabatic upflow conditions using an expanded-bed technology
and cooled recycle. Reactants are converted at moderate well-controlled
temperatures and moderate pressures, maximizing yield and catalyst life.
The main effluents are purified for further applications or recycle. ��
�
More than 90% of the total per pass conversion occurs in the ex-
panded-bed reactor. The effluent then flows to a reactive distillation
system (4), Catacol. This system, operated like a conventional distillation
column, combines catalysis and distillation. The catalytic zones of the
Catacol use fixed-bed arrangements of an inexpensive acidic resin cata-
lyst that is available in bulk quantities and easy to load and unload.
The last part of the unit removes alcohol from the crude raffinate
using a conventional waterwash system (5) and a standard distillation Installation: Over 25 units, including ETBE and TAME, have been licensed.
column (6). Twenty-four units, including four Catacol units, are in operation.
Yields: Ether yields are not only highly dependent on the reactive olefins’ Licensor: Axens.
content and the alcohol’s chemical structure, but also on operating
goals: maximum ether production and/or high final raffinate purity (for
instance, for downstream 1-butene extraction) are achieved.

Economics: Plants and their operations are simple. The same inexpen-
sive (purchased in bulk quantities) and long-lived, non-sophisticated cat-
alysts are used in the main reactor section catalytic region of the Catacol
column, if any.
Ethers—ETBE
Application: The Uhde (Edeleanu) ETBE process combines ethanol and
��
isobutene to produce the high-octane oxygenate ethyl tertiary butyl ��
ether (ETBE).
��
Feeds: C4 cuts from steam cracker and FCC units with isobutene con-
tents ranging from 12% to 30%.

Products: ETBE and other tertiary alkyl ethers are primarily used in gas- ��
��
oline blending as an octane enhancer to improve hydrocarbon com-
bustion efﬁciency. Moreover, blending of ETBE to the gasoline pool will
lower vapor pressure (Rvp). ��

Description: The Uhde (Edeleanu) technology features a two-stage re- ��

actor system of which the first reactor is operated in the recycle mode. ��
With this method, a slight expansion of the catalyst bed is achieved that
ensures very uniform concentration profiles in the reactor and, most
important, avoids hot spot formation. Undesired side reactions, such as
the formation of di-ethyl ether (DEE), are minimized.
The reactor inlet temperature ranges from 50°C at start-of-run to
about 65°C at end-of-run conditions. One important feature of the two- Utility requirements: (C4 feed containing 21% isobutene; per metric
stage system is that the catalyst can be replaced in each reactor sepa- ton of ETBE):
rately, without shutting down the ETBE unit. Steam, LP, kg 110
The catalyst used in this process is a cation-exchange resin and is available Steam, MP, kg 1,000
from several manufacturers. Isobutene conversions of 94% are typical for Electricity, kWh 35
Water, cooling, m3 24
FCC feedstocks. Higher conversions are attainable when processing steam-
cracker C4 cuts that contain isobutene concentrations of about 25%. Installation: The Uhde (Edeleanu) proprietary ETBE process has been
ETBE is recovered as the bottoms product of the distillation unit. The successfully applied in three refineries, converting existing MTBE units.
ethanol-rich C4 distillate is sent to the ethanol recovery section. Water is Two other MTBE plants are in the conversion stage.
used to extract excess ethanol and recycle it back to process. At the top
of the ethanol / water separation column, an ethanol / water azeotrope is Licensor: Uhde GmbH.
recycled to the reactor section. The isobutene-depleted C4 stream may
be sent to a raffinate stripper or to a molsieve-based unit to remove
oxygenates such as DEE, ETBE, ethanol and tert- butanol.
Ethers—MTBE
��
Application: The Uhde (Edeleanu) MTBE process combines methanol ��
and isobutene to produce the high-octane oxygenate—methyl tertiary ��
butyl ether (MTBE).
Feeds: C4-cuts from steam cracker and FCC units with isobutene con-
tents range from 12% to 30%.
Products: MTBE and other tertiary alkyl ethers are primarily used in gas-
oline blending as an octane enhancer to improve hydrocarbon combus-
��
tion efficiency.
Description: The technology features a two-stage reactor system of ��
which the first reactor is operated in the recycle mode. With this meth- ��
od, a slight expansion of the catalyst bed is achieved which ensures very
uniform concentration profiles within the reactor and, most important,
avoids hot spot formation. Undesired side reactions, such as the forma-
tion of dimethyl ether (DME), are minimized.
when the raffinate-stream from the MTBE unit will be used to produce
The reactor inlet temperature ranges from 45°C at start-of-run to
a high-purity butene-1 product.
about 60°C at end-of-run conditions. One important factor of the two-
For a C4 cut containing 22% isobutene, the isobutene conversion
stage system is that the catalyst may be replaced in each reactor sepa-
may exceed 98% at a selectivity for MTBE of 99.5%.
rately, without shutting down the MTBE unit.
The catalyst used in this process is a cation-exchange resin and is Utility requirements, (C4 feed containing 21% isobutene; per metric ton
available from several catalyst manufacturers. Isobutene conversions of of MTBE):
97% are typical for FCC feedstocks. Higher conversions are attainable Steam, MP, kg 100
when processing steam-cracker C4 cuts that contain isobutene concen- Electricity, kWh 35
trations of 25%. Water, cooling, m3 15
MTBE is recovered as the bottoms product of the distillation unit. Steam, LP, kg 900
The methanol-rich C4 distillate is sent to the methanol-recovery section.
Water is used to extract excess methanol and recycle it back to process. Installation: The Uhde (Edeleanu) proprietary MTBE process has been
The isobutene-depleted C4 stream may be sent to a raffinate stripper successfully applied in five refineries. The accumulated licensed capacity
or to a molsieve-based unit to remove other oxygenates such as DME, exceeds 1 MMtpy.
MTBE, methanol and tert-butanol.
Very high isobutene conversion, in excess of 99%, can be achieved Licensor: Uhde GmbH.
through a debutanizer column with structured packings containing ad-
ditional catalyst. This reactive distillation technique is particularly suited
Flue gas denitrification
��
Application: The Topsøe SCR DeNOx process removes NOx from flue ��
gases through reactions with an ammonia-based reducing agent over ��
a specially designed fixed-bed monolithic catalyst. By carefully select-
��
ing the catalyst parameters, channel size and chemical composition, the
process covers a wide range of operating conditions and flue-gas dust ��
��
��
contents and may be applied to practically all types of refinery units in- ��
cluding furnaces, boilers, crackers and FCC units. ��
��
Products: The Topsøe SCR DeNOx converts NOx into inert nitrogen and ��
water vapor. The process may be designed for NOx reductions in excess ��
of 95% and with an ammonia leakage of just a few ppm.
��
Description: The reducing agent such as ammonia or urea, aqueous or
��
pure, is injected into the flue gas stream in stoichiometric proportion to ��
the amount of NOx in the flue gas, controlled by measurement of flue ��
gas flow and NOx concentration. The injection takes place in a grid over
the entire cross-section of the flue-gas duct to ensure a uniform distribu-
tion of NOx and ammonia upstream the SCR catalyst vessel.
The process incorporates Topsøe’s well-proven corrugated mono-
lithic DNX catalyst. DNX is manufactured in small units, which may be Installation: More than 50 refinery units use Topsøe SCR DeNOx catalyst
combined into larger modules to match any requirement in terms of and technology. The applications range from low-dust furnaces to high-
vessel dimensions and pressure drop, and in both horizontal and vertical dust FCC units and temperatures up to 500°C (930°F).
vessel configurations.
The DNX catalyst is based on a fiber-reinforced ceramic carrier, which References: Damgaard, L., B. Widroth and M. Schröter: “Control refin-
gives a unique combination of a high strength and a high micro-poros- ery NOx with SCRs,” Hydrocarbon Processing, November 2004.
ity. The high micro-porosity provides a superior resistance to catalyst
poisons and low weight. The fibers add flexibility to the catalyst so that Licensor: Haldor Topsøe A/S.
it can tolerate a wide range of heating and cooling rates.

Operating conditions: Typical operating conditions range from 300°C to

500°C (570–930°F), up to 3 bar (44 psia) and up to 50 g/Nm3 of dust in
the flue gas.
Flue gas desulfurization—SNOX
Application: The SNOX process treats boiler flue gases from the com-
bustion of high-sulfur fuels, such as heavy residual oil and petroleum
��
��
coke. The SNOX process is a combination of the Topsøe WSA process ��
and the Topsøe SCR DeNOx process. The process removes SO2, SO3 and ��
��
NOx as well as dust. The sulfur is recovered in the form of concentrated
��
commercial-grade sulfuric acid. The SNOX process is distinctly different ��
from most other flue gas desulfurization processes in that its economy
increases with increasing sulfur content in the flue gas.
��
��
Description: Dust is removed from the flue gas by means of an electro- ��
��
��
static precipitator or a bag filter. The flue gas is preheated in a gas/gas ��
��
heat exchanger. Thereafter, it is further heated to approximately 400°C ��
and ammonia is added, before it enters the reactor, where two different ��
��
catalysts are installed. The first catalyst makes the NOx react with ammo-
nia to form N2 and water vapor, and the second catalyst makes the SO2
react with oxygen to form SO3. The second catalyst also removes any
dust traces remaining. During the cooling in the gas/gas heat exchanger,
most of the SO3 reacts with water vapor to form sulfuric acid vapor. The
sulfuric acid vapor is condensed via further cooling in the WSA con-
denser, which is a heat exchanger with vertical glass tubes. • Attractive operating economy
Concentrated commercial-grade sulfuric acid is collected in the • Simple, reliable and flexible process.
bottom of the WSA condenser and is cooled and pumped to storage.
Installation: Four SNOX units have been contracted for cleaning of a to-
Cleaned flue gas leaves the WSA condenser at 100°C and can be sent
tal of more than three million Nm3/ h of flue gas. Additionally, 50 WSA
to the stack without further treatment. The WSA condenser is cooled by
plants have been contracted. These are similar to SNOX plants, only
atmospheric air. The cooling air can be used as preheated combustion
smaller, and some without NOx removal, for other applications than flue
air in the boiler. This process can achieve up to 98% sulfur removal and
gas cleaning.
about 96% NOx removal.
Other features of the SNOX process are: Licensor: Haldor Topsøe A/S.
• No absorbent is applied
• No waste products are produced. Besides dust removed from the
flue gas, the only products are cleaned flue gas and concentrated
commercial-grade sulfuric acid.
• High degree of heat efficiency
• Modest utility consumption
Fluid catalytic cracking
Application: Selective conversion of a wide range of gas oils into high-
value products. Typical feedstocks are virgin or hydrotreated gas oils but �
may also include lube oil extract, coker gas oil and resid.
�
Products: High-octane gasoline, light olefins and distillate. Flexibility of
mode of operation allows for maximizing the most desirable product. �
�
The new commercially proven FCC/ Indmax technology selectively cracks
molecules of different sizes and shapes, thus maximizing light olefins �
(propylene and ethylene) production. �

Description: The Lummus process incorporates an advanced reaction

��
system, high-efﬁciency catalyst stripper and a mechanically robust, � �
single-stage fast ﬂuidized bed regenerator. Oil is injected into the ��

base of the riser via proprietary Micro-Jet feed injection nozzles (1). �
Catalyst and oil vapor flow upwards through a short-contact time,
all-vertical riser (2) where raw oil feedstock is cracked under opti-
mum conditions.
Reaction products exiting the riser are separated from the spent line (7). This arrangement provides the lowest overall unit elevation.
catalyst in a patented, direct-coupled cyclone system (3). Product Catalyst is regenerated by efficient contacting with air for complete
vapors are routed directly to fractionation, thereby eliminating non- combustion of coke. For resid-containing feeds, the optional cata-
selective, post-riser cracking and maintaining the optimum prod- lyst cooler is integrated with the regenerator. The resulting flue gas
uct yield slate. Spent catalyst containing only minute quantities of exits via cyclones (9) to energy recovery/flue gas treating. The hot
hydrocarbon is discharged from the diplegs of the direct-coupled regenerated catalyst is withdrawn via an external withdrawal well
cyclones into the cyclone containment vessel (4). The catalyst flows (10). The well allows independent optimization of catalyst density
down into the stripper containing proprietary modular grid (MG) in the regenerated catalyst standpipe, maximizes slide valve (11)
baffles (5). pressure drop and ensures stable catalyst flow back to the riser feed
Trace hydrocarbons entrained with spent catalyst are removed injection zone.
in the MG stripper using stripping steam. The MG stripper efficient- The catalyst formulation can be tailored to maximize the most
ly removes hydrocarbons at low steam rate. The net stripper vapors desired product. For example, the formulation for maximizing light
are routed to the fractionator via specially designed vents in the olefins (Indmax operation) is a multi-component mixture that pro-
direct-coupled cyclones. Catalyst from the stripper flows down the motes the selective cracking of molecules of different sizes and
spent-catalyst standpipe and through the slide valve (6). The spent shapes to provide very high conversion and yield of light olefins.
catalyst is then transported in dilute phase to the center of the re-
generator (8) through a unique square-bend-spent catalyst transfer
Continued 
Fluid catalytic cracking, continued
Economics:
Investment (basis: 30,000 bpsd including reaction/regeneration
system and product recovery. Excluding offsites, power recovery
and flue gas scrubbing US Gulf Coast 2006.)
$/bpsd (typical) 2,400–3,500
Utilities, typical per bbl fresh feed:
Electricity, kWh 0.8–1.0
Steam, 600 psig (produced) 50–200
Maintenance, % of investment per year 2–3

Installation: Fifteen grassroots units licensed. Twenty-eight units re-

vamped, with ﬁve revamps in design stage.

Licensor: ABB Lummus Global.

Fluid catalytic cracking
��
Application: FLEXICRACKING IIIR converts high-boiling hydrocarbons in-
cluding residues, gas oils, lube extracts and/or deasphalted oils to higher
��
value products. �

Products: Light oleﬁns for gasoline processes and petrochemicals, LPG, �

blend stocks for high-octane gasoline, distillates and fuel oils. �
�
Description: The FLEXICRACKING IIIR technology includes process de- �
sign, hardware details, special mechanical and safety features, control
systems, flue gas processing options and a full range of technical servic- � �
es and support. The reactor (1) incorporates many features to enhance
�
performance, reliability and flexibility, including a riser (2) with patented
high-efficiency close-coupled riser termination (3), enhanced feed injec- �
tion system (4) and efficient stripper design (5). The reactor design and
operation maximizes the selectivity of desired products, such as naphtha
and propylene.
The technology uses an improved catalyst circulation system with
advanced control features, including cold-walled slide valves (6). The Yields: Typical examples:
single vessel regenerator (7) has proprietary process and mechanical fea- Resid feed VGO + lube extracts VGO feed
tures for maximum reliability and efficient air/catalyst distribution and mogas distillate mogas
contacting (8). Either full or partial combustion is used. With increasing operation operation operation
residue processing and the need for additional heat balance control, Feed
partial burn operation with outboard CO combustion is possible, or KBR Gravity, °API 22.9 22.2 25.4
dense phase catalyst cooler technology may be applied. The ExxonMobil Con carbon, wt% 3.9 0.7 0.4
wet gas scrubbing or the ExxonMobil-KBR Cyclofines TSS technologies Quality 80% Atm. Resid 20% Lube Extracts 50% TBP – 794°F
can meet flue gas emission requirements. (Hydrotreated)
Product yields
Naphtha, lv% ff 78.2 40.6 77.6
(C4 / FBP) (C4 / 430°F) (C 4 / 260°F) (C4 / 430°F)
Mid Dist., lv% ff 13.7 49.5 19.2
(IBP / FBP) (430 / 645°F) (260 / 745°F) (430 / 629°F)

Continued 
Fluid catalytic cracking, continued
Installation: More than 70 units with a design capacity of over 2.5-mil-
lion bpd fresh feed.

References: Ladwig, P. K., “Exxon FLEXICRACKING IIIR ﬂuid catalytic

cracking technology,” Handbook of Petroleum Reﬁning Processes, Sec-
ond Ed., R. A. Meyers, Ed., pp. 3.3–3.28.

Licensor: ExxonMobil Research and Engineering Co. and Kellogg Brown

& Root, Inc. (KBR).
Fluid catalytic cracking
Application: Selective conversion of gas oil feedstocks.
Products: High-octane gasoline, distillate and C3– C4 olefins.
Description: Catalytic and selective cracking in a short-contact-time riser
where oil feed is effectively dispersed and vaporized through a propri-
etary feed-injection system. Operation is carried out at a temperature
consistent with targeted yields. The riser temperature profile can be
optimized with the proprietary mixed temperature control (MTC) sys-
tem. Reaction products exit the riser-reactor through a high-efficiency,
close-coupled, proprietary riser termination device RSS (Riser Separator
Stripper). Spent catalyst is pre-stripped followed by an advanced high-
efficiency packed stripper prior to regeneration. The reaction product
vapor may be quenched to give the lowest possible dry gas and maxi-
mum gasoline yield. Final recovery of catalyst particles occurs in cyclones
before the product vapor is transferred to the fractionation section.
Catalyst regeneration is carried out in a single regenerator equipped
with proprietary air and catalyst distribution systems, and may be oper-
ated for either full or partial CO combustion. Heat removal for heavier
feedstocks may be accomplished by using reliable dense-phase cata-
lyst cooler, which has been commercially proven in over 56 units. As Installation: Shaw Stone & Webster and Axens have licensed 27 full-
an alternative to catalyst cooling, this unit can easily be retrofitted to technology units and performed more than 150 revamp projects.
a two-regenerator system in the event that a future resid operation is
desired. Reference: Meyers, R., Handbook of Petroleum Refining Process, Third Ed.
The converter vessels use a cold-wall design that results in minimum
Licensor: Shaw Stone & Webster and Axens, IFP Group Technologies.
capital investment and maximum mechanical reliability and safety. Re-
liable operation is ensured through the use of advanced fluidization
technology combined with a proprietary reaction system. Unit design
is tailored to the refiner’s needs and can include wide turndown flex-
ibility. Available options include power recovery, wasteheat recovery,
flue gas treatment and slurry filtration. Revamps incorporating propri-
etary feed injection and riser termination devices and vapor quench
result in substantial improvements in capacity, yields and feedstock
flexibility within the mechanical limits of the existing unit.
Fluid catalytic cracking
Application: To convert heavy distillates and residues into high-value ��
products, including selective propylene production when required, us- ��
��
ing the Shell FCC Process.
��
Description: In this process, Shell’s high-performance feed nozzle system ��
feeds hydrocarbons to a short contact-time riser; this design ensures ��
good mixing and rapid vaporization into the hot catalyst stream. Crack- ��
��
ing selectivity is enhanced by the feed nozzles and proprietary riser-in- ��
��
��
ternals, which reduce catalyst back mixing while reducing overall riser
��
pressure drop. ��
��
Riser termination design incorporates reliable close-couple cyclones ��
��
��
that provide rapid catalyst/hydrocarbon separation. It minimizes post ��
riser cracking and maximizes desired product yields, with no slurry clean ��
��
up required. Stripping begins inside the first cyclone, followed by a high- ��
capacity baffle structure.
A single-stage partial or full-burn regenerator delivers excellent per-
formance at low cost. Proprietary internals are used at the catalyst inlet
to disperse catalyst, and the catalyst outlet to provide significant catalyst
circulation enhancement. Catalyst coolers can be added for more feed-
stock flexibility.
Cyclone-systems in the reactor and regenerator use a proprietary
design, thus providing reliability, efficiency and robustness. Flue gas
cleanup can be incorporated with Shell’s third-stage separator.
Two FCC design options are available. The Shell 2 Vessel design is
recommended to handle less heavy feeds with mild coking tendencies;
the Shell External Reactor is preferred for heavy feeds with high coking
tendencies. These designs are proven reliability champions due to sim-
plicity of components and incorporation of Shell’s extensive operating
experience.

Installations: Over 30 grassroots units designed/licensed, including 7 to

handle residue feeds, and over 30 units revamped.

Supplier: Shell Global Solutions International B.V.

Fluid catalytic cracking
Application: Selectively convert gas oils and residue feedstocks into
��
higher value products using the FCC/RFCC/PetroFCC process.

Products: Light oleﬁns (for alkylation, polymerization, etheriﬁcation or ��

petrochemicals), LPG, high-octane gasoline, distillates and fuel oils.

Description: UOP’s process uses a side-by-side reactor/regenerator con- ��

figuration and a patented pre-acceleration zone to condition the regen-
erated catalyst. Modern Optimix feed distributors inject the feed into
the riser, which terminates in a vortex separation system (VSS). A high-
efficiency stripper then separates the remaining hydrocarbons from the
catalyst, which is then reactivated in a combustor-style regenerator. With
RxCat technology, a portion of the catalyst that is pre-stripped by the
riser termination device can be recycled back to the riser via a standpipe
and the MxR chamber.
��
The reactor zone features a short-contact-time riser, state-of the-
art riser termination device for quick separation of catalyst and vapor,
with high hydrocarbon containment (VSS/VDS technology) and RxCat
technology, wherein a portion of the pre-stripped (carbonized) catalyst has been licensed for use in four units, two of which are currently in
from the riser termination device is blended with the hotter regenerated construction. The first unit to incorporate RxCat technology has been
catalyst in a proprietary mixing chamber (MxR) for delivery to the riser. operating successfully since the second quarter of 2005.
Unlike other approaches to increasing the catalyst-to-oil ratio, this The combustor-style regenerator burns coke in a fast-fluidized envi-
technology does not affect the total heat balance and, therefore, does ronment completely to CO2 with very low levels of CO. The circulation
not increase coke yield. The reactor temperature can be lowered to re- of hot catalyst from the upper section to the combustor provides added
duce thermal cracking with no negative impact on conversion, thus im- control over the burn-zone temperature and kinetics and enhances ra-
proving product selectivity. The ability to vary the carbonized/regener- dial mixing. Catalyst coolers can be added to new and existing units to
ated catalyst ratio provides considerable flexibility to handle changes in reduce catalyst temperature and increase unit flexibility for commercial
feedstock quality and shortens the time for operating adjustments by operations of feeds up to 6 wt% Conradson carbon. A recent study of
enabling rapid switches between gasoline, olefins or distillate operat- eight different combustor-style regenerators and 15 bubbling-bed re-
ing modes. Since coke yield can be decreased at constant conversion, generators clearly demonstrated that at a given excess oxygen level, less
capacity and reaction severity can be increased, and CO2 emissions re- NOx is emitted from the combustor-style regenerators than other avail-
duced. Furthermore, because the catalyst delivered to the regenerator able technologies.
has a higher coke content, it requires less excess oxygen at a given tem-
perature to sustain the same kinetic combustion rate. RxCat technology Continued 
Fluid catalytic cracking, continued increased conversion, higher selectivity for desired products, the ability
to operate at high catalyst flux and increased capacity, and lower steam
consumption. One application of AF trays resulted in a 0.04 wt% reduc-
For heavier residue feeds, the two-stage regenerator is used. In the
tion in coke, a 14°F (8°C) drop in the regenerator temperature, and a
first stage, upper zone, the bulk of the carbon is burned from the cata-
2.3 LV% boost in conversion. Another FCC unit utilizes these trays to
lyst, forming a mixture of CO and CO2. Catalyst is transferred to the
handle a high catalyst flux rate of over 2-M lb/ft2 min. The installation
second stage, lower zone, where the remaining coke is burned in com-
of AF grids in a small unit increased conversion and gasoline yield by 3.0
plete combustion, producing low levels of carbon on regenerated cata-
LV% and 2.8 LV% respectively, paying back the investment in less than
lyst. A catalyst cooler is located between the stages. This configuration
three months. UOP’s stripper technology has been implemented in more
maximizes oxygen use, requires only one train of cyclones and one flue
than 36 FCC units with a combined on-stream capacity exceeding 1MM
gas stream, which avoids costly multiple flue gas systems and creates a
bpsd.
hydraulically-simple. The two stage regenerator system has processed
feeds up to 8.5 wt% Conradson carbon. Installation: All of UOP’s technology and equipment are commercially
PETROFCC is a customized application using mechanical features proven for both process performance and mechanical reliability. UOP
such as RxCAT technology for recontacting carbonized catalyst, high has been an active designer and licensor of FCC technology since the
severity processing conditions and selected catalyst and additives to early 1940s and has licensed more than 215 FCC, Resid FCC and MSCC
produce high yields of propylene, light olefins and aromatics for petro- process units. More than 150 of these units are operating worldwide. In
chemical applications. addition to applying our technology and skills to new units, UOP is also
UOP’s Advanced Fluidization (AF) spent catalyst stripper internals extensively involved in the revamping of existing units. During the past
provide a family of options (trays, grids and packing) all with state of the 15 years, UOP’s FCC Engineering department has undertaken 40 to 60
art efficiency. Often the optimal selection is dependent on the unique revamp projects or studies per year.
configuration of the unit, site constructability and inspection issues.
Within residence constraints, the new designs can save about 10 feet in Licensor: UOP LLC.
length compared to 1990s units of the same capacity and diameter. The
benefits provided by the AF internals include reduced coke, lower re-
generator temperature, higher catalyst circulation, lower dry gas make,
Fluid catalytic cracking—pretreatment ��
Application: Topsøe’s FCC pretreatment technology is designed to treat ��
a wide variety of feedstocks ranging from gas oils through heavy-vacu- ��

um gas oils and coker streams to resids. This pretreatment process can
��
maximize FCC unit performance. ��

Objectives: The processing objectives range from deep desulfurization ��

for meeting gasoline-sulfur specifications from the FCC products, to de-
��
nitrogenation and metals removal, thus maximizing FCC catalyst activ-
ity. Additional objectives can include Conradson carbon reduction and ��
saturation of polyaromatics to maximize gasoline yields.
��
��
Description: The Topsøe FCC Pretreatment technology combines under- ��
standing of kinetics, high-activity catalysts, state-of-the-art internals and ��
��
engineering skills. The unit can be designed to meet specific process- ��
ing objectives in a cost-effective manner by utilizing the combination of
processing severity and catalyst activity.
Topsøe has experience in revamping moderate- to low-pressure
units for deep desulfurization. Such efforts enable refiners to directly Operating conditions: Typical operating pressures range from 60 to 125
blend gasoline produced from the FCC and meet future low-sulfur (less bar (900 to 1,800 psi), and temperatures from 300°C to 430°C (575°F
than 15 ppm) gasoline specifications. to 800°F).
An additional option is Topsøe’s Aroshift process that maximizes the References: Andonov, G., S. Petrov, D. Stratiev and P. Zeuthen, “MCHC
conversion of polyaromatics which can be equilibrium limited at high mode vs. HDS mode in an FCC unit in relation to Euro IV fuels specifica-
operating temperatures. The Aroshift process increases the FCC conver- tions,” 10th ERTC, Vienna, November 2005.
sion, and the yield of gasoline and C3 /C4 olefins, while reducing the Patel R., H. Moore and B. Hamari, “FCC hydrotreater revamp for low-
amount of light- and heavy-cycle oil. Furthermore, the quality of the sulfur gasoline,” NPRA Annual Meeting, San Antonio, March 2004.
FCC gasoline is improved. Patel, R., P. Zeuthen and M. Schaldemose, “Advanced FCC feed pre-
Topsøe has a wide variety of catalysts for FCC pretreatment service. The treatment technology and catalysts improves FCC profitability,” NPRA
catalyst types cover TK-558 BRIM, a CoMo catalyst with high desulfurization Annual Meeting, San Antonio, March 2002.
activity, and TK-559 BRIM, a NiMo catalyst with hydrodesulfurization and
high hydrodenitrogenation activity. Topsøe offers a wide range of engi- Installations: Four units in the US.
neering scopes from full scoping studies, reactor design packages and Licensor: Haldor Topsøe A/S.
process design packages to engineering design packages.
Gas treating—H2S removal ��
Application: Remove H2S selectively, or remove a group of acidic impuri-
ties (H2S, CO2, COS, CS2 and mercaptans) from a variety of streams, de- ��
pending on the solvent used. FLEXSORB SE technology has been used in ��
refineries, natural gas production facilities and petrochemical operations. �
��
FLEXSORB SE or SE Plus solvent is used on: hydrogenated Claus plant
tail gas to give H2S, ranging down to H2S <10 ppmv; pipeline natural
��
gas to give H2S <0.25 gr/100 scf; or FLEXICOKING low-Btu fuel gas. The
resulting acid gas byproduct stream is rich in H2S. �
Hybrid FLEXSORB SE solvent is used to selectively remove H2S, as
well as organic sulfur impurities commonly found in natural gas.

Description: A typical amine system flow scheme is used. The feed gas
contacts the treating solvent in the absorber (1). The resulting rich sol-
vent bottom stream is heated and sent to the regenerator (2). Regen-
erator heat is supplied by any suitable heat source. Lean solvent from
the regenerator is sent through rich/lean solvent exchangers and coolers
before returning to the absorber.
Reference: Garrison, J., et al., “Keyspan Energy Canada Rimbey acid gas
FLEXSORB SE solvent is an aqueous solution of a hindered amine.
enrichment with FLEXSORB SE Plus technology,” 2002 Laurance Reid
FLEXSORB SE Plus solvent is an enhanced aqueous solution, which has
Gas Conditioning Conference, Norman, Oklahoma.
improved H2S regenerability yielding <10 vppm H2S in the treated gas.
Adams-Smith, J., et al., Chevron USA Production Company, “Carter
Hybrid FLEXSORB SE solvent is a hybrid solution containing FLEXSORB SE
Creek Gas Plant FLEXSORB tail gas treating unit,” 2002 GPA Annual
amine, a physical solvent and water.
Meeting, Dallas.
Economics: Lower investment and energy requirements based primarily Connock, L., et al., “High recovery tail gas treating,” Sulphur, No.
on requiring 30% to 50% lower solution circulation rates, compared to 296, November/ December 2004.
conventional amines. Fedich, R., et al., “Selective H2S Removal,” Hydrocarbon Engineer-
ing, May 2004.
Installations: Total gases treated by FLEXSORB solvents are about 2 bil- Fedich, R. B., et al., “Solvent changeover benefits,” Hydrocarbon
lion scfd and the total sulfur recovery is about 900 long tpd. Engineering, Vol. 10, No. 5, May 2005.
FLEXSORB SE—31 plants operating, three in design “Gas Processes 2006,” Hydrocarbon Processing, January 2006.
FLEXSORB SE Plus—19 plants operating, nine in design
Hybrid FLEXSORB SE—two plants operating, three in design Licensor: ExxonMobil Research and Engineering Co.
Over 60 plants operating or in design.
Gasification
Application: The Shell Gasification Process (SGP) converts the heaviest
��
residual liquid hydrocarbon streams with high-sulfur and metals content
��
into a clean synthesis gas and valuable metal oxides. Sulfur (S) is re- ��
��
��
moved by normal gas treating processes and sold as elemental S.
The process converts residual streams with virtually zero value as ��
��
fuel-blending components into valuable, clean gas and byproducts. ��
This gas can be used to generate power in gas turbines and for making
H2 by the well-known shift and PSA technology. It is one of the few ul-
timate, environmentally acceptable solutions for residual hydrocarbon
streams. ��
��

Products: Synthesis gas (CO+H2), sulfur and metal oxides. ��

Process description: Liquid hydrocarbon feedstock (from very light

such as natural gas to very heavy such as vacuum flashed cracked
residue, VFCR and ashphalt) is fed into a reactor, and gasified with
pure O2 and steam. The net reaction is exothermic and produces a
gas primarily containing CO and H2. Depending on the final syngas
application, operating pressures, ranging from atmospheric up to 65
A related process—the Shell Coal Gasification Process (SCGP)—gas-
bar, can easily be accommodated. SGP uses refractory-lined reactors
ifies solids such as coal or petroleum coke. The reactor is different, but
that are fitted with both burners and a heat-recovery-steam generator,
main process layout and work-up are similar.
designed to produce high-pressure steam—over 100 bar (about 2.5
tons per ton feedstock). Gases leaving the steam generator are at a Installation: Over the past 40 years, more than 150 SGP units have been
temperature approaching the steam temperature; thus, further heat installed that convert residue feedstock into synthesis gas for chemical
recovery occurs in an economizer. applications. The latest, flagship installation is in the Shell Pernis refinery
Soot (unconverted carbon) and ash are removed from the raw gas near Rotterdam, The Netherlands. This highly complex refinery depends
by a two-stage waterwash. After the final scrubbing, the gas is virtually on the SGP process for its H2 supply. Similar projects are underway in
particulate-free; it is then routed to a selective-acid-gas-removal system. Canada and Italy.
Net water from the scrubber section is routed to the soot ash removal The Demkolec Power plant at Buggenum, The Netherlands pro-
unit (SARU) to filter out soot and ash from the slurry. By controlled oxi- duces 250 Mwe based on the SCGP process. The Shell middle distillate
dation of the filtercake, the ash components are recovered as valuable synthesis plant in Bintulu, Malaysia, uses SGP to convert 100 million scfd
oxides—principally vanadium pentoxide. The (clean) filtrate is returned of natural gas into synthesis gas that is used for petrochemical applica-
to the scrubber. tions.
Continued 
Gasification, continued
Reference: “Shell Gasification Process,” Conference Defining the Fu-
ture, Bahrain, June 1–2, 2004.
“Shell Gasification Process for Upgrading Gdansk Refinery,” The
6th European Gasification Conference IChemE, Brighton, May 10–12,
2004.
“Overview of Shell Global Solutions Worldwide Gasification Devel-
opments,” 2003 Gasification Technologies Conference, San Francisco,
Oct. 12–15, 2003.

Licensor: Shell Global Solutions International B.V.

Gasoline desulfurization
Application: Convert high-sulfur gasoline streams into a low-sulfur gas- ��
oline blendstock while minimizing octane loss, yield loss and operating ��
cost using S Zorb sulfur removal technology. ��

��
Products: Load-sulfur blending stock for gasoline motor fuels.
Description: Gasoline from the ﬂuid catalytic cracker unit is combined
with a small hydrogen stream and heated. Vaporized gasoline is injected

��
into the ﬂuid-bed reactor (1), where the proprietary sorbent removes ��
��
sulfur from the feed. A disengaging zone in the reactor removes sus- ��
pended sorbent from the vapor, which exits the reactor to be cooled. ��

��
Regeneration: The sorbent (catalyst) is continuously withdrawn from the ��
reactor and transferred to the regenerator section (2), where the sulfur ��
��
is removed as SO2 and sent to a sulfur-recovery unit. The cleansed sor-
bent is reconditioned and returned to the reactor. The rate of sorbent
circulation is controlled to help maintain the desired sulfur concentra-
tion in the product.

Economics:
Typical operating conditions: Results:
Temperature, °F 750 – 825 C5+ yield, vol% of feed >100%
Pressure, psig 100 – 500 Lights yield, wt% of feed < 0.1
Space velocity, whsv 4–8 (R+M) loss
Hydrogen purity, % 70 – 99 2 <0.3
Total H2 usage, scf / bbl 40 – 60 Operating cost, ¢/gal* 0.9
Case study premises: * Includes utilities, 4% per year maintenance and sorbent costs.
25,000 - bpd feed
Installation: Forty-three sites licensed as of 1Q 2004.
775 - ppm feed sulfur
25 - ppm product sulfur ( 97% removal ) Licensor: ConocoPhillips.
No cat gasoline splitter
Gasoline desulfurization, ultra-deep
Application: Ultra-deep desulfurization of FCC gasoline with minimal
��
octane penalty using Prime-G+ process.
��
Description: FCC debutanizer bottoms are fed directly to a first reactor ��
��
wherein, under mild conditions, diolefins are selectively hydrogenated
and mercaptans are converted to heavier sulfur species. The selective
��
hydrogenation reactor effluent is then usually split to produce an LCN ��
(light cat naphtha) cut and an HCN (heavy cat naphtha). ��
The LCN stream is mercaptans-free with a low-sulfur and diolefin ��
concentration, enabling further processing in an etherification or al- ��
��
kylation unit. The HCN then enters the main Prime-G+ section where
it undergoes in a dual catalyst reactor system; a deep HDS with very
limited olefins saturation and no aromatics losses produces an ultra- ��
low-sulfur gasoline. ��

The process provides ﬂexibility to advantageously co-process other

sulfur-containing naphthas such as light coker naphtha, steam cracker
naphtha or light straight-run naphtha.
Industrial results:
Full-range FCC Gasoline, Prime-G+ OATS process: In addition to Prime-G+ TAME and TAEE etherification
40°C–220°C Feed Product technology, the OATS (olefins alkylation of thiophenic sulfur) process,
Sulfur, ppm 2,100 50* initially developed by BP, is also exclusively offered for license by Axens
(RON + MON) / 2 87.5 86.5 for ultra-low-sulfur gasoline production.
 (RON + MON) / 2 1.0 Reference: “Prime-G+: From pilot to start-up of world’s first commercial
% HDS 97.6 10 ppm FCC gasoline desulfurization process,” NPRA Annual Meeting,
 30 ppm pool sulfur after blending March 17–19, 2002, San Antonio.
Pool sulfur specifications as low as less than 10 ppm are attained
with the Prime-G+ process in two units in Germany. Licensor: Axens.
Economics:
Investment: Grassroots ISBL cost, $/bpsd 600–800
Installation: Currently, 126 units have been licensed for a total capacity
of 3.3 million bpsd. Seventy Prime-G+ units are already in operation,
producing ultra-low-sulfur gasoline.
Gasoline desulfurization, ultra-deep
Application: Reduce sulfur in gasoline to less than 10 ppm by ��
hydrodesulfurization followed by cracking and isomerization to recover ��
octane with the OCTGAIN process.
��
Description: The basic flow scheme of the OCTGAIN process is similar to ��
��
that of a conventional naphtha hydrotreater. Feed and recycle hydrogen ��
��
mix is preheated in feed/effluent exchangers and a fired heater then in- ��
��
troduced into a fixed-bed reactor. Over the first catalyst bed, the sulfur
in the feed is converted to hydrogen sulfide (H2S) with near complete ��
olefin saturation. In the second bed, over a different catalyst, octane is ��
recovered by cracking and isomerization reactions. The reactor effluent ��
��
is cooled and the liquid product separated from the recycle gas using
��
high- and low-temperature separators. ��
The vapor from the separators is combined with makeup gas, com-
pressed and recycled. The liquid from the separators is sent to the prod-
uct stripper where the light ends are recovered overhead and desulfur-
ized naphtha from the bottoms. The product sulfur level can be as low
as 5 ppm. The OCTGAIN process can be retrofitted into existing refinery
hydrotreating units. The design and operation permit the desired level
of octane recovery and yields.
EMRE has an alliance with Kellogg Brown & Root (KBR) to provide
this technology to refiners.

Yields: Yield depends on feed oleﬁns and desired product octane.

Installations: Commercial experience with two operating units.
Reference: Halbert, T., et al., “Technology Options For Meeting Low-Sul-
fur Mogas Targets,” NPRA Annual Meeting, March 2000.

Licensor: ExxonMobil Research and Engineering Co.

Gasoline desulfurization, ultra-deep
Application: Reduce sulfur in FCC gasoline to levels as low as <10 wppm
by selective hydrotreating to maximize octane retention with the SCAN- ��
ﬁning technology. ��

Description: The feed is mixed with hydrogen, heated with reactor efflu- ��
ent exchange and passed through a pretreat reactor for diolefin satura- ��
��
tion. After further heat exchange with reactor effluent and preheat using ��
��
a utility, the hydrocarbon/hydrogen mixture enters the main reaction sec- ��
��
tion which features ExxonMobil Research and Engineering Co. (EMRE)
proprietary selective catalyst systems. In this section of the plant, sulfur ��
is removed in the form of H2S under tailored process conditions, which ��
��
strongly favor hydrodesulfurization while minimizing olefin saturation. ��
The feed may be full-range, intermediate or heavy FCC-naphtha ��
fraction. Other sulfur-containing streams such as light-coker naphtha, ��
steam cracker or light straight-run naphthas can also be processed with
FCC naphthas. SCANfining technology can be retrofitted to existing
units such as naphtha or diesel hydrotreaters and reformers. SCANfining
technology also features ExxonMobil’s proprietary reactor internals such
as Automatic Bed Bypass Technology for onstream mitigation of reactor
plugging/pressure drop buildup. References: Sapre, A.V., et al., “Case History: Desulfurization of FCC
For high-sulfur feeds and/or very low-sulfur product, with low levels naphtha,” Hydrocarbon Processing, February 2004.
of product mercaptans variations in the plant design from SCANfining I Ellis, E. S., et al., “Meeting the Low Sulfur Mogas Challenge,” World
Process to the SCANfining II Process for greater HDS selectivity, or addi- Refining Association Third European Fuels Conference, March 2002.
tion of a ZEROMER process step for mercaptan conversion, or addition
of an EXOMER process unit for mercaptan extraction. Licensor: ExxonMobil Research and Engineering Co.
EMRE has an alliance with Kellogg Brown & Root (KBR) to pro-
vide SCANfining technology to refiners and an alliance with Merichem
Chemicals & Refinery Services LLC to provide EXOMER technology to
refiners.

Yields: Yield of C5+ liquid product is typically over 100 LV%.

Installation: Thirty-seven units under design, construction or operation
having combined capacity of over 1.1 million bpsd.
H2S and SWS gas conversion
Application: The ATS process recovers H2S and NH3 in amine regenera-
tor offgas and sour water stripper gas (SWS gas) as a 60% aqueous ��
solution of ATS – ammonium thiosulfate (NH 4 )2S 2 O 3, which is the stan-
��
dard commercial speciﬁcation. The ATS process can be combined with ��
��
a Claus unit; thus increasing processing capacity while obtaining a total ��
sulfur recovery of > 99.95%. The ATS process can also handle SWS gas
��
alone without ammonia import while the S / N balance is adjusted by
exchanging H 2S surplus and deﬁcit with a Claus unit. ��
ATS is increasingly used as a fertilizer (12- 0 - 0 -26S) for direct appli- ��

cation and as a component in liquid fertilizer formulations.

Description: Amine regenerator off gas is combusted in a burner / waste ��

��
heat boiler. The resulting SO2 is absorbed with ammonia in a two-stage
��
absorber to form ammonium hydrogen sulﬁte (AHS). NH3 and H 2 S con-
tained in the SWS gas plus imported ammonia (if required) is reacted
with the AHS solution in the ATS reactor. The ATS product is withdrawn
as a 60% aqueous solution that meets all commercial speciﬁcations for
usage as a fertilizer. Unreacted H 2S is returned to the H 2S burner.
Except for the H 2S burner / waste heat boiler, all process steps occur
in the liquid phase at moderate temperatures and neutral pressure. The
AHS absorber and ATS reactor systems are chilled with cooling water.
More than 99.95% of the sulfur and practically 100% of the am-
monia contained in the feed gas streams are recovered. Typical emission
values are:
SOx <100 ppmv
NOx <50 ppmv
H 2S <1 ppmv
NH3 <20 ppmv

Installation: One Topsøe 30,000 mtpy ATS plant is operating in Northern

Europe.

Licensor: Haldor Topsøe A/S.

H2S removal ��

Application: LO-CAT removes H2S from gas streams and produces el-
��
emental sulfur. LO-CAT units are in service treating reﬁnery fuel gas,
hydrodesulfurization offgas, sour-water-stripper gas, amine acid gas,
��
claus tail gas and sulfur tank vent gas. Sulfur capacities are typically less ��
than 25 ltpd down to several pounds per day. Key beneﬁts of operation � �

are high (99.9%) H2S removal efficiency, and flexible operation, with vir- ��
tually 100% turndown capability of H2S composition and total gas flow.
��
Sulfur is recovered as a slurry, filter cake or high-purity molten sulfur.
The sulfur cake is increasingly being used in agriculture, but can also be ��
deposited in a nonhazardous landfill.
��
��
Description: The conventional configuration is used to process combus-
tible gas and product gas streams. Sour gas contacts the dilute, propri-
etary, iron chelate catalyst solution in an absorber (1), where the H2S is
absorbed and oxidized to solid sulfur. Sweet gas leaves the absorber for
use by the refinery. The reduced catalyst solution returns to the oxidizer restrictions on type of gas to be treated; however, some contaminants,
(2), where sparged air reoxidizes the catalyst solution. The catalyst solu- such as SO2, may increase operating costs.
tion is returned to the absorber. Continuous regeneration of the catalyst Installations: Presently, 160 licensed units are in operation with four
solution allows for very low chemical operating costs. units under construction.
In the patented autocirculation configuration, the absorber (1) and
oxidizer (2) are combined in one vessel, but separated internally by baf- Reference: Nagl, G., W. Rouleau and J. Watson,, “Consider optimized
fles. Sparging of the sour gas and regeneration air into the specially Iron-Redox processes to remove sulfur,” Hydrocarbon Processing, Janu-
designed baffle system creates a series of “gas lift” pumps, eliminating ary 2003, pp. 53–57.
the external circulation pumps. This configuration is ideally suited for
treating amine acid gas and sour-water-stripper gas streams. Licensor: Gas Technology Products, a division of Merichem Chemical &
In both configurations, sulfur is concentrated in the oxidizer cone Refinery Services LLC.
and sent to a sulfur filter, which can produce filter cake as high as 85%
sulfur. If desired, the filter cake can be further washed and melted to
produce pure molten sulfur.

Operating conditions: Operating pressures range from vacuum condi-

tions to 1,000 psi. Operating temperatures range from 40°F to 140°F.
Hydrogen sulfide concentrations range from a few ppm to 100%. Sul-
fur loadings range from a few pounds per day to 25+ tons per day. No
H2S removal
Applications: Sulfur-Rite is a solid-bed scavenger for removal of H2S from
aerobic and anaerobic gas streams. Suitable applications are generally
sulfur loads below 200 lb/d sulfur, and/or remote refinery locations. Sul-
fur vents, loading and unloading facilities, or backup insurance for other
refinery sulfur-removal systems are examples.
��
The spent media is nonpyrophoric, and is suitable for disposition in
nonhazardous landfills. ��
��
Description: Single-bed (shown) or dual “lead-lag” configurations are ��
possible. Sour gas is saturated prior to entering media bed. Gas enters ��

vessel top, flows over media where H2S is removed and reacted. Sweet ��
gas exits the bottom of vessel. In the single-vessel configuration, when
the H2S level exceeds the level allowed, the vessel must be bypassed,
media removed through the lower manway, fresh media installed and
vessel returned to service.
For continuous operation, a dual “lead-lag” configuration is desir-
able. The two vessels operate in series, with one vessel in the lead posi-
tion, the other in the lag position. When the H2S level at the outlet of
the lead vessel equals the inlet H2S level (the media is completely spent),
the gas flow is changed and the vessels reverse rolls, so that the “lag”
vessel becomes the “lead” vessel. The vessel with the spent media is
bypassed. The media is replaced, and the vessel with fresh media is re-
turned to service in the “lag” position.

Operating conditions: Gas streams up to 400°F can be treated. Gas

streams should be at least 50% water saturated.

Installations: Sixteen units installed.

Licensor: Gas Technology Products, a division of Merichem Chemical &
Refinery Services LLC.
the removal of spent solution and the addition of fresh solution without
H2S removal shutting down. This arrangement also results in the optimum utilization
of the solution.
Application: ELIMINATOR technology consisting of a full line of ELIMI- • Packed tower—Sour gas is contacted with circulating solution of
NATOR products removes hydrogen sulfide (H2S) and light mercaptans ELIMINATOR in a counter, packed-bed scrubber.
from a variety of process gas streams in a safe, efficient, environmentally
friendly and easy to operate manner. Products: A full line of ELIMINATOR products can treat any type of gas
streams.
Description: The ELIMINATOR technology is extremely versatile, and its
performance is not sensitive to operating pressure. In properly designed Economics: Operating costs are very favorable for removing less than
systems, H2S concentrations of less than 1 ppm can easily be achieved 250 kg/d of H2S.
on a continuous basis.
A number of different treatment methodologies may be used to Installation: Fifteen units in operation.
treat sour gas streams.
Licensor: Gas Technology Products, a division of Merichem Chemicals &
• Line injection—ELIMINATOR can be sprayed directly into a gas
Refinery Services LLC.
stream with removal of the spent product in a downstream knockout
pot.
• Sparge tower—Sour gas is bubbled up through a static volume of
ELIMINATOR. A lead-lag vessel arrangement can be installed to allow for
Hydroconversion—VGO and DAO
Application: An ebullated-bed process H-OilDC is used for hydrocon- ��
��
��
version (hydrocracking and hydrotreating) of heavy vacuum gasoil and ��
��
DAO having high Conradson carbon residue and metal contents and ��
low asphaltene content. It is best suited for high severity operations and ��
applications requiring long run lengths. ��
��
��
��
Description: The flow diagram includes integrated mid-distillate hydro- ��

treating for an ultra-low-sulfur-diesel product. The typical battery limits ��

scheme includes oil- and hydrogen-ﬁred heaters, an advanced design ��
hot high-pressure separator and ebullating pump recycle system, a re- ��
��
cycle gas scrubber and product separation and fractionation. ��
��
Catalyst in the reactor is replaced periodically without shutdown ��
��
and, for cases of feeds with low metal contents, the catalyst can be re- ��
generated onsite to reduce catalyst consumption. ��
Various catalysts are available as a function of the feedstock and
the required objectives. An H-OilDC unit can operate for four-year run
lengths at constant catalyst activity with conversion in the 20-80% range
in once-through mode and to more than 95% in recycle mode with up
to 99% hydrodesulfurization.
Economics: Basis—2005 US Gulf Coast
Investment in $ per bpsd 2,500 – 4,000
Operating conditions: Utilities, per bbl of feed
Temperature, °F 750– 820 Fuel, 103 Btu 60
Hydrogen partial pressure, psi 600 –1,500 Power, kWh 3
LHSV, hr –1 0.5 –3.0 Catalyst makeup, lb 0.01– 0.3
Conversion, wt% 20 –80 in once-through mode
Installation: The H-OilDC process has two references, one in operation
Example: VGO + DAO feed: a blend of heavy VGO and C5 DAO con- and one under construction, with a total cumulative capacity of 139,900
taining close to 100 ppm metals is processed at 80% conversion at an bpsd. This technology has been commercially demonstrated based on
overall desulfurization rate of over 96%. the ebullated bed reactor making a total of nine references for residue
and VGO hydroconversion.

Licensor: Axens.
Hydrocracking
��
Application: Upgrade vacuum gas oil alone or blended with various
feedstocks (light-cycle oil, deasphalted oil, visbreaker or coker gasoil). ��

Products: Jet fuel, diesel, very-low-sulfur fuel oil, extra-quality FCC ��
��
feed with limited or no FCC gasoline post-treatment or high VI lube
base stocks. ��
�
� � � �
��
Description: This process uses a refining catalyst usually followed by
an amorphous and/or zeolite-type hydrocracking catalyst. Main features ��
of this process are:
• High tolerance toward feedstock nitrogen
• High selectivity toward middle distillates ��
• High activity of the zeolite, allowing for 3–4 year cycle lengths and ��
��
products with low aromatics content until end of cycle.
Three different process arrangements are available: single-step/
once-through; single-step/total conversion with liquid recycle; and two-
step hydrocracking. The process consists of: reaction section (1, 2), gas
separator (3), stripper (4) and product fractionator (5). Economics:
Investment: (Basis: 40,000-bpsd unit, once-through, 90% con-
Product quality: Typical for HVGO (50/50 Arabian light/heavy):
version, battery limits, erected, engineering fees included, 2000
Feed, Jet Gulf Coast), $ per bpsd 2,500 –3,500
HVGO fuel Diesel Utilities, typical per bbl feed:
Sp. gr. 0.932 0.800 0.826 Fuel oil, kg 5.3
TBP cut point, °C 405– 565 140 –225 225 –360 Electricity, kWh 6.9
Sulfur, ppm 31,700 <10 <10 Water, cooling, m 3 0.64
Nitrogen, ppm 853 <5 <5 Steam, MP balance
Metals, ppm <2 – –
Cetane index – – 62 Installation: More than 50 references, cumulative capacity exceeding
Flash pt., °C –  40 125 1 million bpsd, conversions up to 99%.
Smoke pt., mm, EOR – 26–28 – Licensor: Axens.
Aromatics, vol%, EOR – < 12 <8
Viscosity @ 38°C, cSt 110 – 5.3
PAH, wt%, EOR <2
Hydrocracking ��
Application: Convert a wide variety of feedstocks including vacuum
deep-cut gas oil, coker gas oils, de-asphalted oil (DAO), and FCC cy-
cle oils into high-quality, low-sulfur fuels using ExxonMobil Research ��
and Engineering Company’s (EMRE) moderate pressure hydrocracking ��
(MPHC) process. ��
��
��
Products: Products include a wide range of high-quality, low-sulfur dis- ��
��
tillates and blending stocks including LPG, high-octane gasoline, high- ��
quality reformer naphtha. Unconverted bottoms product from the MPHC ��
unit is very low in sulfur and is an excellent feedstock for fluid catalytic ��
cracking (FCC), lube-oil basestock production, steam cracking and low- ��
sulfur fuel oil.
��
Description: The process uses a multiple catalyst system in multi-bed ��
reactor(s) that incorporates proprietary advanced quench and redistri-
bution internals (Spider Vortex). Heavy hydrocarbons and recycle gas
are preheated and contact the catalyst in the trickle-phase fixed-bed
reactor(s). Reactor effluent is flashed in high- and low-temperature
separators. An amine scrubber removes H2S from the recycle gas be-
fore it gets compressed and re-circulated back to the unit. An opti-
mized, low cost stripper/fractionator arrangement is used for product Yields:
recovery. Naphtha, wt% 4 10 10
When higher-quality distillates are required, the addition of a low- Kero/jet, wt% 6 10 10
cost, highly integrated distillate post-treating unit (PTU) can be incor- Diesel, wt% 22 26 27
porated in the design to meet or exceed high-pressure hydrocracking LSGO (FCC feed), wt% 65 50 50
product quality at lower capital cost and hydrogen consumption H2 consumption, wt% 1.0 –1.5 1.3 –1.8 1.5 – 2.0
Product quality:
Operating conditions and yields: Typical operating conditions on a Mid-
Kero sulfur, wppm 20 – 200 20 – 200 20 – 200
dle East VGO for a once-through MPHC operation are shown:
Kero smoke Pt, mm 13 – 18 15 – 20 17 – 22
Operation conditions: Diesel sulfur, wppm 30 – 500 30 – 300 30 – 200
Configuration MPHC MPHC MPHC Diesel cetane no. 45 – 50 47 – 52 50 – 55
Nominal conversion, % 35 50 50
H2 pressure, psig 800 800 1,250
Continued 
Hydrocracking, continued
Utilities, per bbl of feed:
Electric power, kW 4.1 7.2
Fuel (absorbed), Btu 67,100 69,600
Steam, MP (export), lb (15.9) (21.1)
Water, cooling, gal 101 178
Wash water, gal 1.5 2.2
Lean amine, gal 36.1 36.1

EMRE’s MPHC process is equally amenable to revamp or grassroots

applications. EMRE has an alliance with Kellogg Brown & Root (KBR) to
provide MPHC technology to reﬁners.

Economics: Investment $/ bpsd 2,000 – 3,000

Installation: Four operating units; two in construction.

Licensor: ExxonMobil Research and Engineering Co.
Hydrocracking
Application: Topsøe’s hydrocracking process can be used to convert ��
straight run vacuum gas oils and heavy cracked gas oils to high quality ��
“sulfur-free” naphtha, kerosene, diesel, and FCC feed, meeting current
��
and future regulatory requirements. In addition, high VI lube stocks and ��
petrochemical feedstock can be produced to increase the reﬁnery’s prof- ��
��
itability. ��
��
Product: By proper selection of operating conditions, process conﬁgura-
��
tion, and catalysts, the Topsøe hydrocracking process can be designed ��
��
for high conversion to produce high smoke point kerosine and high ��
��
cetane diesel. The process can also be designed for lower conversion/ ��
��
upgrade mode to produce low sulfur FCC feed with the optimum hy-
��
drogen uptake or high VI (>145) lube stock. The FCC gasoline produced
��
from a Topsøe hydrocracking unit does not require post-treatment for ��
sulfur removal. ��

Description: Topsøe’s hydrocracking process uses well proven co-current

downflow fixed bed reactors with state - of - the - art reactor internals and
catalysts. The process uses recycle hydrogen and can be configured in
partial conversion once-through feed mode or with recycle of partially
converted oil to obtain 100% conversion to diesel and lighter prod-
ucts. Topsøe’s zeolitic and amorphous hydrocracking catalysts have been
proven in several commercial hydrocrackers.

Operating conditions: Typical operating pressure and temperatures

range from 55 to 170 bar (800 to 2500 psig) and 340 to 420°C (645 to
780°F).

Installations: One operating licensed hydrocracking unit. Topsøe hydro-

cracking catalysts have been supplied to eight hydrocrackers.

Licensor: Haldor Topsøe A/S.

Hydrocracking ��
��
Application: Process to upgrade ultra-heavy oil to high-quality distillates ��
��
using the KOBELCO SPH (Slurry-Phase Hydrocracking) process. ��
��
��
��
Description: In the KOBELCO SPH, ultra-heavy oil is hydrocracked via a
slurry-bed reactor (1). In the hydrocracking (HC) step, a dispersed natural � �
��
��
limonite ore (-FeOOH) is used as the catalyst. The hydrocracked products �

are sent to an inline hydrotreating (HT) step and a solid-liquid separation ��

(SLS) step. The HT step is designed to apply economically temperatures ��

��
and pressures from the hydrocracking reactions. It consists of two-stage
��
fixed-bed reactors (2) filled with conventional hydrotreating catalysts.
��
The SLS step applies vacuum flashing (3) and toluene-insoluble (TI) ��
removal (4). The TI (coke) and limonite catalyst that adsorb heavy met- �

als are removed from the residual. The solid-free fraction from the top ��

of the TI-remover is combined with the heavy fraction from the vacuum ��

ﬂasher and recycled to the hydrocracking reactor.

Typical operating conditions, product yields and properties for atmo-
spheric topped bitumen (ATB) from Athabasca oil sands as the feed are:
Economics: For a 55,000-bpsd facility installed with an existing upgrader
Feed properties: API gravity, 6.7°; 524°C+ residue, 48.5 wt%; sulfur, 5.0 in northern Alberta, Canada. The construction cost is estimated at
wt%; nitrogen, 0.44 wt%; carbon residue, 14.3 wt%; TI, 0.61 wt%. US$409 million (2004 basis).
Utilities (per bbl fresh feed):
Operating conditions: Fuel, thousand kcal 30
Hydrocracking: 450°C, 10MPa, 1 hr residence time, 0.5 wt% Electricity, kWh 19
as Fe catalyst loading: Water, cooling, m 3 2.2
1st stage hydrotreating: 350°C, 10MPa, 1 hr –1 LHSV
2nd stage hydrotreating: 350°C, 10MPa, 0.5 hr –1 LHSV Installations: No commercial units; a 3-bpd demonstration unit has been
operated for over 1,500 hours.
Product yields and properties:
Products Naphtha Diesel fuel VGO Residue Reference: Okui, T., M. Yasumuro, M.Tamura, T. Shigehisa, and S. Yui,
Cut range, °C C5 –177 177–343 343–524 524+ “Convert oil sands into distillate cost-effectively,” Hydrocarbon Process-
Total yields, vol% on ATB 19.7 52.0 30.8 3.0 ing, January 2006, pp.79-85.
First stage HT / Second stage HT
Licensor: Kobe Steel Ltd.
Sulfur, ppm (wt) 27/1 141/2 1600/–
Nitrogen, ppm (wt) 9/1 91/1 1550/–
Total hydrogen consumption: 2.8 wt% on ATB (314 Nm3/m3 ATB).
Hydrocracking
Application: To convert heavy VGO and other low-cost cracked and ex- ��
tracted feedstocks into high-value, high-quality products, such as low-
sulfur diesel, jet fuel, high-octane light gasoline and reformer feed via ��
��
the Shell Hydrocracking Process. Unconverted or recycle oil are prime ��
feeds for secondary processing in FCCUs, lube base oil plants and eth- ��
��
ylene crackers. � ��
��
��
Description: Heavy feed hydrocarbons are preheated with reactor efflu- �
ent (1). Fresh hydrogen is combined with recycle gas from the cold high- ��
pressure separator, preheated with reactor effluent, and then heated in ��
� ��
a single-phase furnace. Reactants pass via trickle flow through multi-bed ��
��
reactor(s) containing proprietary pre-treat, cracking and post-treat cata- ��
lysts (2). Interbed ultra-flat quench internals and high dispersion nozzle ��
��
trays combine excellent quench, mixing and liquid flow distribution at
the top of each catalyst bed while maximizing reactor volume utiliza-
tion. After cooling by feed streams, reactor effluent enters a separator
system. Hot effluent is routed to fractionation (3).
Two-stage, series flow and single-stage unit design configurations
are available including a single-reactor, stacked-bed design suitable for
capacities up to 10,000 tpd in partial or full-conversion modes. The
catalyst systems are carefully tailored for the desired product slate and
catalyst cycle length.

Installations: Over 30 new and revamp designs installed or under de-

sign. Revamps have been implemented in own or other licensors’ de-
signs usually to debottleneck and increase feed heaviness.

Supplier: Shell Global Solutions International B.V.

Hydrocracking
Application: Convert a wide variety of feedstocks into lower-molecular-
��
weight products using the Unicracking and HyCycle Unicracking process. ��
��
Feed: Feedstocks include atmospheric gasoil, vacuum gasoil, FCC/RCC cycle
oil, coker gasoil, deasphalted oil and naphtha for production of LPG. �

Products: Processing objectives include production of gasoline, jet fuel, �

� ��
diesel fuel, lube stocks, ethylene-plant feedstock, high-quality FCC feed-
stock and LPG.
�
Description: Feed and hydrogen are contacted with catalysts, which in- ��
duce desulfurization, denitrogenation and hydrocracking. Catalysts are
��
based upon both amorphous and molecular-sieve containing supports. � ��
Process objectives determine catalyst selection for a speciﬁc unit. Prod- ��
uct from the reactor section is condensed, separated from hydrogen-rich
gas and fractionated into desired products. Unconverted oil is recycled
or used as lube stock, FCC feedstock or ethylene-plant feedstock.
Yields: Example:
FCC cycle Vacuum Fluid coker Economics: Example:
Feed type oil blend gasoil gasoil Investment, $ per bpsd capacity 2,000–4,000
Gravity, °API 27.8 22.7 8.4 Utilities, typical per bbl feed:
Boiling, 10%, °F 481 690 640 Fuel, 103 Btu 70–120
End pt., °F 674 1,015 1,100 Electricity, kWh 7–10
Sulfur, wt% 0.54 2.4 4.57
Installation: Selected for 161 commercial units, including several converted
Nitrogen, wt% 0.024 0.08 0.269
from competing technologies. Total capacity exceeds 3.6 million bpsd.
Principal products Gasoline Jet Diesel FCC feed
Yields, vol% of feed Licensor: UOP LLC.
Butanes 16.0 6.3 3.8 5.2
Light gasoline 33.0 12.9 7.9 8.8
Heavy naphtha 75.0 11.0 9.4 31.8
Jet fuel 89.0
Diesel fuel 94.1 33.8
600°F + gas oil 35.0
H2 consump., scf/bbl 2,150 1,860 1,550 2,500
Hydrocracking
Application: Convert naphthas, AGO, VGO, DAO, cracked oils from FCC �
units, delayed cokers and visbreakers, and intermediate products from
residue hydroprocessing units using the Chevron Lummus Global ISO- � �
CRACKING Process. ��
�

Products: Lighter, high-quality, more valuable products: LPG, gasoline, ��

catalytic reformer feed, jet fuel, kerosene, diesel and feeds for FCC, eth- ��
��
ylene cracker or lube oil units. ��
��
��
Description: A broad range of both amorphous/zeolite and zeolitic �
��
catalysts, including noble-metal zeolitic catalysts, are used to tailor the � �
ISOCRACKING Process exactly to the reﬁner’s objectives. In general, ��
��
the process involves a staged reactor system with an initial stage of ��

hydrotreating and partially hydrocracking the feed, and a subsequent ��

stage continuing the conversion or product upgrade process in a more � ��
��
favorable environment. ��
Feeds can be introduced in between stages using Chevron Lummus
Global patented split-feed injection technology, or effluent flow paths can
be arranged to best utilize hydrogen and minimize quench-gas requirements Yields: Typical from various feeds:
using proprietary SSRS (single-stage reaction sequenced) technology. Feed VGO VGO VGO VGO
Most modern large-capacity flow schemes involving heavy sour gas Gravity, API 24.1 24.1 24.1 21.3
oils require two reactors (1, 4) and one high-pressure separation system TBP range, °F 700–1,100 700–1,100 700–1,100 700–1,100
(2) with an optional recycle gas scrubber (5) and one recycle-gas com- Nitrogen, wppm 2,500 2,500 2,500 900
pressor (8). The low-pressure separators (3), product stripper (6) and Sulfur, wt % 1.9 1.9 1.9 2.5
fractionator (7) provide the flexibility to fractionate products either in Mode Max. Diesel Max. Jet Max. Mid- Max. Mid-
between reaction stages or at the tail-end, depending on desired prod- Distillate Distillate
uct slate and selectivity requirements. + Lubes
Single-stage options are used in once-through mode typically for Yields, vol %
mild hydrocracking or when a significant quantity of unconverted oil is Naphtha 22.8 30.8 14.0 18
required for FCC, lubes, or ethylene units. The single-stage recycle op- Jet/kerosine – 79.7 22.0 50
tion is used for lower capacity units when economical. The reactors use Diesel 85.5 – 73.0 35
patented internals technology called ISOMIX for near-flawless mixing UCO – – – 10
and redistribution.

Continued 
Hydrocracking, continued
Feed VGO VGO VGO VGO
Product quality
Kerosine smoke, mm 29–32 29–32 29–32
Diesel cetane number 58–64 58–64 58–64
UCO BMCI 6–8
UCO Waxy V.I. 143–145
UCO Dewaxed V.I. 131–133

Economics: ISBL total installed cost of 35,000-BPSD unit at 100% con-

version to middle distillates using Middle Eastern VGO feed (USGC, mid-
2006 basis): $150 million.
Process fuel (absorbed), MMBtu/hr 180
Electricity, MW 10
CW, gpm 2,500
Steam (export at 150 psig), M lb/hr 22

Installation: More than 60 units worldwide with over one million-bpsd

total capacity.

Licensor: Chevron Lummus Global LLC.

Hydrocracking—residue ��

Application: H-OilRC is an ebullated-bed process for hydrocracking at- ��

��
mospheric or vacuum residue. It is the ideal solution for feedstocks hav- ��
��
ing high metal, CCR and asphaltene contents. The process can have two ��
��
different objectives: at high conversion, to produce stable products; or, ��
at moderate conversion, to produce a synthetic crude oil. ��
��
��
Description: The ﬂow diagram illustrates a typical H-OilRC unit that in-
��
cludes oil and hydrogen ﬁred heaters, an optional inter stage separa- ��
��
tor, an internal recycle cup providing feed to the ebullating pump, high ��
pressure separators, recycle gas scrubber and product separation and ��
fractionation (not required for synthetic crude oil production).
��
Catalyst is replaced periodically in the reactor, without shutdown. ��
Different catalysts are available as a function of the feedstock and ��
��
the required objectives. An H-OilRC unit can operate for three-year run ��
��
lengths at constant catalyst activity with conversion in the 50 – 80%
range and hydrodesulfurization as high as 85%.

Operating conditions:
Temperature, °F 770– 820 Economics: Basis 2005 US Gulf Coast
Hydrogen partial pressure, psi 1,600 –1,950 Investment in $ per bpsd 4,500 – 6,500
LHSV, hr –1 0.25– 0.6 Utilities, per bbl of feed
Conversion, wt% 50 – 80 Fuel, 103 Btu 70
Power, kWh 11
Examples: Ural VR feed: a 540°C+ cut from Ural crude is processed at Catalyst makeup, lb 0.2– 0.8
66% conversion to obtain a stable fuel oil containing less than 1%wt
sulfur, 25% diesel and 30% VGO. The diesel cut is further hydrotreated Installation: There are seven H-OilRC units in operation with a total capac-
to meet ULSD speciﬁcations using an integrated Prime-D unit. Arab Me- ity of 300,000 bpsd. Two additional references for H-OilDC, the ebullated
dium VR feed: a vacuum residue from a blend 70% Arab Light-30% bed technology for VGO and DAO, add another 139,900 bpsd.
Arab Heavy containing 5.5wt% sulfur is processed at above 75% con- Licensor: Axens.
version to obtain a stable fuel oil with 2wt% sulfur.
Hydrocracking
Application: Desulfurization, demetalization, CCR reduction and hydro-
cracking of atmospheric and vacuum resids using the LC-FINING Pro- ��

cess. ��
��
Products: Full range of high-quality distillates. Residual products can ��
be used as fuel oil, synthetic crude or feedstock for a resid FCC, coker, �
visbreaker or solvent deasphalter.
�
�
Description: Fresh hydrocarbon liquid feed is mixed with hydrogen and ��
reacted within an expanded catalyst bed (1) maintained in turbulence by ��
liquid upflow to achieve efficient isothermal operation. Product quality is ��
maintained constant and at a high level by intermittent catalyst addition
and withdrawal. Reactor products flow to a high-pressure separator (2), ��
low-pressure separator (3) and product fractionator (4). Recycle hydro-
gen is separated (5) and purified (6).
Process features include onstream catalyst addition and withdrawal.
Recovering and purifying the recycled H2 at low pressure rather than at high
pressure can reduce capital cost and allows design at lower gas rates.

Operating conditions:
Reactor temperature, °F 725 – 840
Atm. resid Vac. resid
Reactor pressure, psig 1,400 – 3,500 Feed
Gravity, °API 12.40 4.73 4.73 4.73
H2 partial pressure, psig 1,000 – 2,700
Sulfur, wt % 3.90 4.97 4.97 4.97
LSHV 0.1 to 0.6
Ni / V, ppmw 18 /65 39 /142 39 /142 39 /142
Conversion, % 40 – 97+ Conversion, vol% 45 60 75 95
Desulfurization, % 60 – 90 (1,022°F+)
Demetalization, % 50 – 98 Products, vol%
CCR reduction, % 35 – 80 C4 1.11 2.35 3.57 5.53
C5–350°F 6.89 12.60 18.25 23.86
Yields: For Arabian heavy/Arabian light blends: 350 –700°F (650°F) (15.24) 30.62 42.65 64.81
700 (650°F) –1,022°F (55.27) 21.46 19.32 11.92
1,022°F+ 25.33 40.00 25.00 5.0
C5+, °API / wt% S 23.70 / 0.54 22.5 / 0.71 26.6 / 0.66 33.3 / 0.33

Continued 
Hydrocracking, continued
Economics:
Investment, estimated (US Gulf Coast, 2006)
Size, bpsd fresh feed 92,000 49,000
$/bpsd typical fresh feed 3,000 5,000 5,800 7,200
Utilities, per bbl fresh feed
Fuel ﬁred, 103 Btu 56.1 62.8 69.8 88.6
Electricity, kWh 8.4 13.9 16.5 22.9
Steam (export), lb 35.5 69.2 97.0 97.7
Water, cooling, gal. 64.2 163 164 248

Installation: Six LC-FINING units are in operation, and two LC-FINING

Units are in engineering.

Licensor: Chevron Lummus Global LLC.

Hydrodearomatization
��
Application: Topsøe’s two-stage hydrodesulfurization hydrodearomati- ��
��
zation (HDS/HDA) process is designed to produce low-aromatics distil- ��
��
late products. This process enables refiners to meet the new, stringent ��
��
��
standards for environmentally friendly fuels. ��
��
Products: Ultra-low sulfur, ultra-low nitrogen, low-aromatics diesel, ker- ��
��
osine and solvents (ultra-low aromatics). ��
��
Description: The process consists of four sections: initial hydrotreating,
intermediate stripping, final hydrotreating and product stripping. The ��
initial hydrotreating step, or the “first stage” of the two-stage reaction ��
��
process, is similar to conventional Topsøe hydrotreating, using a Topsøe ��
high-activity base metal catalyst such as TK-575 BRIM to perform deep ��
desulfurization and deep denitrification of the distillate feed. Liquid ef- ��
��
fluent from this first stage is sent to an intermediate stripping section, in ��
which H2S and ammonia are removed using steam or recycle hydrogen.
Stripped distillate is sent to the final hydrotreating reactor, or the “second
stage.” In this reactor, distillate feed undergoes saturation of aromatics
using a Topsøe noble metal catalyst, either TK-907/TK-911 or TK-915, a
high-activity dearomatization catalyst. Finally, the desulfurized, dearoma- Operating conditions: Typical operating pressures range from 20 to 60
tized distillate product is steam stripped in the product stripping column barg (300 to 900 psig), and typical operating temperatures range from
to remove H2S, dissolved gases and a small amount of naphtha formed. 320°C to 400°C (600°F to 750°F) in the first stage reactor, and from
Like the conventional Topsøe hydrotreating process, the HDS/HDA 260°C to 330°C (500°F to 625°F) in the second stage reactor. An ex-
process uses Topsøe’s graded bed loading and high-efficiency patented ample of the Topsøe HDS/HDA treatment of a heavy straight-run gas oil
reactor internals to provide optimum reactor performance and catalyst feed is shown below:
use leading to the longest possible catalyst cycle lengths. Topsøe’s high Feed Product
efficiency internals have a low sensitivity to unlevelness and are designed Specific gravity 0.86 0.83
to ensure the most effective mixing of liquid and vapor streams and max- Sulfur, ppmw 3,000 1
imum utilization of catalyst. These internals are effective at high liquid Nitrogen, ppmw 400 <1
loadings, thereby enabling high turndown ratios. Topsøe’s graded-bed Total aromatics, wt% 30 <10
technology and the use of shape-optimized inert topping and catalysts Cetane index, D-976 49 57
minimize the build-up of pressure drop, thereby enabling longer catalyst
cycle length.
Continued 
Hydrodearomatization, continued
References: Cooper, Hannerup and Søgaard-Andersen, “Reduction of
aromatics in diesel,” Oil and Gas, September 1994.
de la Fuente, E., P. Christensen, and M. Johansen, “Options for meet-
ing EU year 2005 fuel specifications,” 4th ERTC, Paris, November 1999.

Installation: A total of seven units.

Licensor: Haldor Topsøe A/S.
Hydrofinishing
Application: Deeply saturate single- and multiple-ring aromatics in base-
oil feedstocks. The product will have very low-aromatics content, very ��
high-oxidation stability and high thermal stability.
Description: ISOFINISHING catalysts hydrogenate aromatics at relative-
ly low reaction temperatures. They are especially effective in complete
polyaromatics saturation—a reaction that is normally equilibrium limited. �

Typical feedstocks are the effluent from a dewaxing reactor, effluent from
hydrated feeds or solvent-dewaxed feedstocks. The products are highly
stabilized base-oil, technical-grade white oil or food-grade white oil. ��
As shown in the simplified flow diagram, feedstocks are mixed with ��
�
recycle hydrogen and fresh makeup hydrogen, heated and charged to a ��
�
reactor containing ISOFINISHING Catalyst (1). Effluent from the finishing ��

reactor is ﬂashed in high-pressure and low-pressure separators (2, 3). ��

A very small amount of light products are recovered in a fractionation
system (4).
Yields: For a typical feedstock, such as dewaxing reactor effluent, the
yield can be >99%. The chemical-hydrogen consumption is usually very
low, less than ~10 Nm3/m3 oil.
Economics:
Investment: For a stand-alone ISOFINISHING Unit, the ISBL capital is
about 3,000 –5,000 $/bpsd, depending on the pressure
level and size.
Utilities: Typical per bbl feed:
Power, kW 2.6
Fuel, kcal 3.4 x 103
Installation: Twenty-five units are in various stages of operation, con-
struction or design.
Reference: NPRA Annual Meeting, March 2004, San Antonio, Paper
AM-04-68.
Licensor: Chevron Lummus Global LLC.
Hydrofinishing/hydrotreating
Application: Process to produce finished lube-base oils and special oils. ��
��
��
Feeds: Dewaxed solvent or hydrogen-refined lube stocks or raw vacuum ��
distillates for lubricating oils ranging from spindle oil to machine oil and ��
bright stock.
��
Products: Finished lube oils (base grades or intermediate lube oils) and ��
��
special oils with specified color, thermal and oxidation stability. ��
��
Description: Feedstock is fed together with make-up and recycle hydro-
gen over a fixed-bed catalyst at moderate temperature and pressure.
The treated oil is separated from unreacted hydrogen, which is recycled. ��
Very high yields product are obtained. ��
For lube-oil hydrofinishing, the catalytic hydrogenation process is
operated at medium hydrogen pressure, moderate temperature and low ��
��
hydrogen consumption. The catalyst is easily regenerated with steam
and air.
Operating pressures for hydrogen-finishing processes range from
25 to 80 bar. The higher-pressure range enables greater flexibility with
regard to base-stock source and product qualities. Oil color and thermal
stability depend on treating severity. Hydrogen consumption depends
on the feed stock and desired product quality.
Utility requirements (typical, Middle East Crude), units per m3 of feed:
Electricity, kWh 15
Steam, MP, kg 25
Steam, LP, kg 45
Fuel oil, kg 3
Water, cooling, m3 10
Installation: Numerous installations using the Uhde (Edeleanu) propri-
etary technology are in operation worldwide. The most recent reference
is a complete lube-oil production facility licensed to the state of Turk-
menistan.
Licensor: Uhde GmbH.
Hydrogen
Application: Production of hydrogen ( H2 ) from hydrocarbon (HC) feed- ��
��
stocks, by steam reforming. ��
��
��
Feedstocks: Ranging from natural gas to heavy naphtha as well as po- ��
��
��
��
tential refinery offgases. Many recent refinery hydrogen plants have
multiple feedstock flexibility, either in terms of back-up or alternative
or mixed feed. Automatic feedstock change-over has also successfully ��
��
��
been applied by TECHNIP in several modern plants with multiple feed- ��

��
stock ﬂexibility. ��

��
��
��
��
Description: The generic flowsheet consists of feed pretreatment, pre- ��
��
reforming (optional), steam-HC reforming, shift conversion and hydro- ��
��
gen purification by pressure swing adsorption (PSA). However, it is often ��
tailored to satisfy specific requirements.
��
Feed pretreatment normally involves removal of sulfur, chlorine and ��
other catalyst poisons after preheating to 350°C to 400°C.
The treated feed gas mixed with process steam is reformed in a fired
reformer (with adiadatic pre-reformer upstream, if used) after neces-
sary superheating. The net reforming reactions are strongly endother-
mic. Heat is supplied by combusting PSA purge gas, supplemented by
drogen recovery and generation, and recuperative (post-)reforming also
makeup fuel in multiple burners in a top-fired furnace.
for capacity retrofits.
Reforming severity is optimized for each specific case. Waste heat
from reformed gas is recovered through steam generation before the Installations: TECHNIP has been involved in over 240 hydrogen plants
water-gas shift conversion. Most of the carbon monoxide is further con- worldwide, covering a wide range of capacities. Most installations are
verted to hydrogen. Process condensate resulting from heat recovery for refinery application with basic features for high reliability and opti-
and cooling is separated and generally reused in the steam system after mized cost.
necessary treatment. The entire steam generation is usually on natural
circulation, which adds to higher reliability. The gas flows to the PSA unit Licensor: Technip.
that provides high-purity hydrogen product (up to < 1ppm CO) at near
inlet pressures.
Typical specific energy consumption based on feed + fuel – export
steam ranges between 3.0 and 3.5 Gcal / KNm 3 ( 330 – 370 Btu / scf ) LHV,
depending upon feedstock, plant capacity, optimization criteria and
steam-export requirements. Recent advances include integration of hy-
Hydrogen
Application: Production of hydrogen for refinery applications (e.g., ��
hydrotreating and hydrocracking) as well as for petrochemical and other ��
industrial uses. ��
��
Feed: Natural gas, refinery offgases, LPG, naphtha or mixtures thereof. ��
� �
��
Product: High-purity hydrogen (typically >99.9%), CO, CO2, HP steam
and/or electricity may be produced as separate creditable byproduct.
��
Description: The plant generally comprises four process units. The feed
is desulfurized (1), mixed with steam and converted to synthesis gas in a ��
steam reformer (2) over a nickel-containing catalyst at 20 – 40 bar pres- ��
��
sure and outlet temperatures of 800 – 900°C. ��

The Uhde steam reformer features a well-proven, top-ﬁred design ��

with tubes made of centrifugally cast alloy steel and a unique propri-
etary “cold” outlet manifold system for enhanced reliability. A further
specialty is Uhde’s bi-sectional steam system for the environment-friend-
ly full recovery of process condensate and production of contaminant-
free high-pressure export steam (3) with a proven process gas cooler
design. The Uhde steam reformer concept also includes a modularized Economics: Depending on the individual plant concept, the typical con-
shop-tested convection bank to maximize plant quality and minimize sumption figure for natural gas based plants (feed + fuel – steam) may
construction risks. be lower than 3.13 Gcal /1,000 Nm3 (333 MMBtu/ MMscf) or 3.09 (329)
The final process units are the adiabatic carbon monoxide (CO) shift with prereforming.
(4) and the pressure swing adsorption unit (5) to obtain high-purity hy-
drogen. Process options include feed evaporation, adiabatic feed pre- Installation: Uhde has recently been awarded several designs for hydro-
reforming and/or HT/LT shift to process, e.g., heavier feeds and/or opti- gen plants. These include a 150,000 Nm³/ h (134 MMscfd) plant for Shell
mize feed/fuel consumption and steam production. in Canada, a 91,000 Nm³/ h (81 MMscfd) plant for Bayernoil in Germany
Uhde’s design allows combining maximized process heat recovery and and a 100,000 Nm³/ h (89 MMscfd) plant for SINCOR C.A. in Venezuela.
optimized energy efficiency with operational safety and reliability. Uhde In addition, Uhde is cooperating with Caloric, Germany, in executing a
usually offers tailor-made designs based on either its own or the custom- contract from Shell for a 10,000 Nm3/h (9 MMscfd) plant in Argentina.
er’s design standards. The hydrogen plant is often fully integrated into the This marks the first success in the partnership of Uhde and Caloric in
refinery, particularly with respect to steam production and the usage of the field of smaller-sized hydrogen plants. During 2006, Uhde will also
refinery waste gases. Furthermore, Uhde has wide experience in the con- startup Europe’s largest hydrogen plant for Neste Oil Corp. of Finland,
struction of highly reliable large-scale reformers for hydrogen capacities of formerly Fortum Oil, with a capacity of 155,000 Nm³/ h (139 MMscfd).
up to 220,000 Nm3/h (197 MMscfd) in single-train configurations. Continued 
Hydrogen, continued
References: Ruthardt, K., K. R. Radtke and J. Larsen, “Hydrogen trends,”
Hydrocarbon Engineering, November 2005, pp. 41– 45.
Michel, M., “Design and Engineering Experience with Large-Scale
Hydrogen Plants,” Oil Gas European Magazine, Vol. 30 (2004) No. 2 in:
Erdöl Erdgas Kohle Vol. 120 (2004) No. 6, pp. OG 85– 88.

Licensor: Uhde GmbH.

Hydrogenation
Application: The CDHydro process is used to selectively hydrogenate diole-
��
ﬁns in the top section of a hydrocarbon distillation column. Additional ap-
plications—including mercaptan removal, hydroisomerization and hydro- ��
��
genation of oleﬁns and aromatics are also available.

Description: The patented CDHydro process combines fractionation

with hydrogenation. Proprietary devices containing catalyst are installed
�
in the fractionation column’s top section (1). Hydrogen is introduced ��
�
beneath the catalyst zone. Fractionation carries light components into
��
the catalyst zone where the reaction with hydrogen occurs. Fraction- ��
��
ation also sends heavy materials to the bottom. This prevents foulants
and heavy catalyst poisons in the feed from contacting the catalyst. In
addition, clean hydrogenated reﬂux continuously washes the catalyst ��

zone. These factors combine to give a long catalyst life. Additionally, ��
mercaptans can react with diolefins to make heavy, thermally-stable sul-
fides. The sulfides are fractionated to the bottoms product. This can
eliminate the need for a separate mercaptan removal step. The distillate
product is ideal feedstock for alkylation or etherification processes.
The heat of reaction evaporates liquid, and the resulting vapor is
condensed in the overhead condenser (2) to provide additional reflux.
Economics: Fixed-bed hydrogenation requires a distillation column fol-
lowed by a fixed-bed hydrogenation unit. The CDHydro process elimi-
The natural temperature profile in the fractionation column results in a
nates the fixed-bed unit by incorporating catalyst in the column. When
virtually isothermal catalyst bed rather than the temperature increase
a new distillation column is used, capital cost of the column is only 5%
typical of conventional reactors.
to 20% more than for a standard column depending on the CDHydro
The CDHydro process can operate at much lower pressure than
application. Elimination of the fixed-bed reactor and stripper can reduce
conventional processes. Pressures for the CDHydro process are typically
capital cost by as much as 50%.
set by the fractionation requirements. Additionally, the elimination of a
separate hydrogenation reactor and hydrogen stripper offers significant Installation: Forty-five CDHydro units are in commercial operation for
capital cost reduction relative to conventional technologies. C4, C5, C6 and benzene hydrogenation applications. Nineteen units
Feeding the CDHydro process with reformate and light-straight run have been in operation for more than five years and total commer-
for benzene saturation provides the refiner with increased flexibility to cial operating time now exceeds 100 years for CDHydro technologies.
produce low-benzene gasoline. Isomerization of the resulting C5 / C6 Twelve additional units are currently in engineering / construction.
overhead stream provides higher octane and yield due to reduced ben-
zene and C7+ content compared to typical isomerization feedstocks. Licensor: CDTECH.
Hydrogen—HTCR and HTCR twin
plants
��
Application: Produce hydrogen from hydrocarbon feedstocks such as:
natural gas, LPG, naphtha, refinery offgases, etc., using the Haldor Top-
søe Convective Reformer (HTCR). Plant capacities range from approxi-
mately 5,000 Nm3/ h to 25,000+ Nm3/ h (5 MM scfd to 25+ MMscfd) and ��
��
hydrogen purity from about 99.5 – 99.999+%. This is achieved without
��
any steam export.

Description: The HTCR-based hydrogen plant can be tailor-made to suit

the customer’s needs with respect to feedstock ﬂexibility. A typical plant ��

comprises feedstock desulfurization, pre-reforming, HTCR reforming, ��

��
shift reaction and pressure swing adsorption (PSA) purification to obtain ��
product-grade hydrogen. PSA offgases are used as fuel in the HTCR. Ex-
cess heat in the plant is efficiently used for process heating and process
steam generation.
A unique feature of the HTCR is the high thermal efficiency. Product
gas and flue gas are cooled by providing heat to the reforming reac-
tion to about 600°C (1,100°F). The high thermal efficiency is utilized
to design energy-efficient hydrogen plants without the need for steam
export. In larger plants, the reforming section consists of two HTCR re-
formers operating in parallel.

Economics: HTCR-based hydrogen plants provide the customer with

a low-investment cost and low operating expenses for hydrogen pro-
duction. The plant is supplied as a skid-mounted unit providing a short
installation time. These plants provide high operating ﬂexibility, reliabil-
ity and safety. Fully automated operation, startup and shutdown allow
minimum operator attendance. A net energy efﬁciency of about 3.4
Gcal / 1,000 Nm3 hydrogen (361 MMBtu /scf H2) is achieved depending
on size and feedstock.

Installations: Twenty-eight licensed units.

Licensor: Haldor Topsøe A/S.
Hydrogen—HTER-p
Application: Topsøe’s proprietary and patented HTER-p (Haldor Top-
��
søe Exchange Reformer— parallel installation) technology is a revamp
option for production increase in a steam-reforming-based hydrogen ��
plant. The technology allows hydrogen capacity increases of more than ��
25%. This option is especially advantageous because the signiﬁcant ca-
pacity expansion is possible with minimal impact on the existing tubular
reformer, which usually is the plant bottleneck.

Description: The HTER-p is installed in parallel with the tubular steam ��
methane reformer (SMR) and fed independently with desulfurized ��
feed taken upstream the reformer section. This enables individual ad- ��
��
justment of feedrate and steam- and process steam-to-carbon ratio to
obtain the desired conversion. The hydrocarbon feed is reformed over
a catalyst bed installed in the HTER-p. Process efﬂuent from the SMR ��
is transferred to the HTER-p and mixed internally with the product gas
from the HTER-p catalyst. The process gas supplies the required heat
for the reforming reaction in the tubes of the HTER-p. Thus, no addi-
tional ﬁring is required for the reforming reactions in the HTER-p.

Economics: An HTER-p offers a compact and cost-effective hydrogen

capacity expansion. The investment cost is as low as 60% of that for a
new hydrogen plant. Energy consumption increases only slightly. For a
25% capacity increase, the net energy consumption is 3.13 Gcal / 1,000
Nm3 H 2 (333 MM Btu / scf H 2 ).

References: Dybkjær, I., and S. W. Madsen, “Novel Revamp Solutions

for Increased Hydrogen Demands,” Eighth European Reﬁning Technol-
ogy Conference, November 17–19, 2003, London, UK

Licensor: Haldor Topsøe A/S.

Hydrogen—Methanol-to-Shift
Application: Topsøe’s proprietary Methanol-to-Shift technology is a re-
��
vamp option for hydrogen production increase for a reforming-based ��
��
hydrogen plant. This technology can raise hydrogen production ca- ��
��
pacity by more than 25%. The capacity expansion is flexible and can ��
be changed in very short time; the technology is suitable for capacity
peak shaving and offers the refiner higher feedstock and product slate
flexibility. ��
��
��
��
Description: Additional hydrogen is produced by reforming of methanol
over Topsøe’s novel dual-function catalyst—LK-510. When installed in the ��
existing CO shift converter and fed simultaneously with methanol and re- ��
��
formed gas, the LK-510 catalyst promotes both the conversion of CO with
��
steam to H2 and CO2 and the reforming of methanol to H2 and CO2.
Methanol from a day tank is pumped to a steam-heated evaporator ��
��
and fed as vapor to the existing CO shift converter, now loaded with
the LK-510 catalyst. In most cases, it will be necessary to revamp the
PSA unit for the additional capacity and to check the equipment down-
stream of the CO shift converter and modify as required.

Economics: The Methanol-to-Shift revamp technology is a low-invest-

ment option for hydrogen capacity increase and is rapid to install. The
total investment cost is less than 40% of that of a new hydrogen plant.
Methanol consumption is approximately 0.54 kg / Nm3 hydrogen (0.03
lb/scf H2).

References: Dybkjær, I., and S. W. Madsen, “Novel Revamp Solutions

for Increased Hydrogen Demands,” Eighth European Reﬁning Technol-
ogy Conference, November 17–19, 2003, London, UK.

Licensor: Haldor Topsøe A/S.

Hydrogen—recovery
Application: To recover and purify hydrogen or to reject hydrogen from
reﬁnery, petrochemical or gas processing streams using a PRISM mem- ��

brane. Reﬁnery streams include hydrotreating or hydrocracking purge,

catalytic reformer offgas, ﬂuid catalytic cracker offgas or fuel gas. Pet-
rochemical process streams include ammonia synthesis purge, methanol ��
�
synthesis purge or ethylene offgas. Synthesis gas includes those gener-
� �
ated from steam reforming or partial oxidation. ��
��
Product: Typical hydrogen (H2) product purity is 90 – 98% and, in some
cases, 99.9%. Product purity is dependent upon feed purity, available ��
differential partial pressure and desired H2 recovery level. Typical H2 re-
��
covery is 80–95% or more.
The hydrocarbon-rich nonpermeate product is returned at nearly
the same pressure as the feed gas for use as fuel gas, or in the case of
synthesis gas applications, as a carbon monoxide (CO) enriched feed to
oxo-alcohol, organic acid, or Fisher-Tropsch synthesis.

Description: Typical PRISM membrane systems consist of a pretreatment Economics: Economic benefits are derived from high-product recoveries
(1) section to remove entrained liquids and preheat feed before gas en- and purities, from high reliability and low capital cost. Additional ben-
ters the membrane separators (2). Various membrane separator con- efits include relative ease of operation with minimal maintenance. Also,
figurations are possible to optimize purity and recovery, and operating systems are expandable and adaptable to changing requirements.
and capital costs such as adding a second stage membrane separator
Installations: Over 270 PRISM H2 membrane systems have been com-
(3). Pretreatment options include water scrubbing to recover ammonia
missioned or are in design. These systems include over 54 systems in re-
from ammonia synthesis purge stream.
finery applications, 124 in ammonia synthesis purge and 30 in synthesis
Membrane separators are compact bundles of hollow fibers contained in
gas applications.
a coded pressure vessel. The pressurized feed enters the vessel and flows on
the outside of the fibers (shell side). Hydrogen selectively permeates through Licensor: Air Products and Chemicals, Inc.
the membrane to the inside of the hollow fibers (tube side), which is at lower
pressure. PRISM membrane separators’ key benefits include resistance to wa-
ter exposure, particulates and low feed to nonpermeate pressure drop.
Membrane systems consist of a pre-assembled skid unit with pres-
sure vessels, interconnecting piping, and instrumentation and are fac-
tory tested for ease of installation and commissioning.
Hydrogen—steam reforming ��
��
Application: Production of hydrogen for refinery hydrotreating and hydro- ��
cracking or other refinery, petrochemical and other uses.
�
Feedstock: Light hydrocarbons such as natural gas, refinery fuel gas, LPG/bu- �
tane mixed pentanes and light naphtha.
�
Product: High-purity hydrogen (99.9+%) at any required pressure. �
Description: The feed is heated in the feed preheater and passed through
the hydrotreater (1). The hydrotreater converts sulfur compounds to H2S and �
saturates any unsaturated hydrocarbons in the feed. The gas is then sent to
the desulfurizers (2). These adsorb the H2S from the gas.
The desulfurized feed gas is mixed with steam and superheated in the
feed preheat coil. The feed mixture then passes through catalyst-filled tubes in ��
the reformer (3). In the presence of nickel catalyst, the feed reacts with steam
to produce hydrogen and carbon oxides. Heat for the endothermic reform-
ing reaction is provided by carefully controlled external firing in the reformer. Economics: Typical utilities per Mscf of hydrogen production based on a
Combustion air preheat is used, if applicable to limit export steam. natural gas feedstock and maximum export steam:
Gas leaving the reformer is cooled by the process steam generator (4). Gas Feed and fuel, MM Btu LHV 0.44
is then fed to the shift converter (5), which contains a bed of copper-promoted Export steam, lb 75
iron-chromium catalyst. This converts CO and water vapor to additional H2 Boiler feedwater, lb 115
and CO2. Shift converter effluent gas is cooled, condensate is separated and Power, kW 0.5
the gas is sent to a PSA hydrogen purification system (6). Water, cooling, gal 10
The PSA system operates on a repeated cycle having two basic steps: ad-
sorption and regeneration. PSA offgas is sent to the reformer, where it provides Installations: Over 175 plants worldwide — ranging in size from less than
most of the fuel requirement. Hydrogen from the PSA unit is sent off plot. A 1 MMscfd to over 120 MMscfd capacities. Plant designs for capacities
small hydrogen stream is recycled to the feed of the plant for hydrotreating. from 1 to 200 MMscfd.
The thermal efficiency of the plant is optimized by recovery of heat
from the reformer flue gas stream and from the reformer effluent process Supplier: CB&I Howe Baker.
gas stream. This energy is utilized to preheat reformer feed gas and generate
steam for reforming and export. The process design is customized for each
application depending on project economics and export steam demand.
Hydrogen—steam reforming
Application: Manufacture hydrogen for hydrotreating, hydrocracking or
other refinery or chemical use. �
��
��
Feedstock: Light saturated hydrocarbons: refinery gas or natural gas,
��
LPG or light naphtha.
� �
Products: Typical purity 99.99%; pressure 300 psig, with steam or CO2
as byproducts.

Description: Hydrogen is produced by steam reforming of hydrocarbons ��

with purification by pressure swing adsorption (PSA). Feed is heated (1) ��
�
and then hydrogenated (2) over a cobalt-molybdenum catalyst bed fol-
lowed by purification (3) with zinc oxide to remove sulfur. The purified ��
feed is mixed with steam and preheated further, then reformed over ��
nickel catalyst in the tubes of the reforming furnace (1).
Foster Wheeler’s Terrace Wall reformer combines high efficiency with
ease of operation and reliability. Depending on size or site requirements,
Foster Wheeler can also provide a down-fired furnace. Combustion air
preheat can be used to reduce fuel consumption and steam export. Air Steam
Pre-reforming can be used upstream of the reformer if a mixture preheat generation
of naphtha and light feeds will be used, or if steam export must be Natural gas, feed + fuel, MMBtu/hr 780 885
minimized. The syngas from the reformer is cooled by generating steam, Export steam at 600 psig/700ºF, lb/hr 35,000 130,000
then reacted in the shift converter (4) where CO reacts with steam to Boiler feedwater, lb/hr 70,000 170,000
form additional H2 and CO2. Electricity, kW 670 170
In the PSA section (5), impurities are removed by solid adsorbent, Water, cooling, 18ºF rise, gpm 350 350
and the adsorbent beds are regenerated by depressurizing. Purge gas Installations: Over 100 plants, ranging from less than 1 MMscfd to 95
from the PSA section, containing CO2, CH4, CO and some H2, is used as MMscfd in a single train, with numerous multi-train installations.
fuel in the reforming furnace. Heat recovery from reformer flue gas may
be via combustion air preheat or additional steam generation. Variations Reference: Handbook of Petroleum Refining Processes, Third Ed., Mc-
include a scrubbing system to recover CO2. Graw-Hill, 2003, pp 6.3–6.33.

Economics: Licensor: Foster Wheeler.

Investment: 10–100 MMscfd, First Q 2005, USGC $10– 60 million
Utilities, 50 MMscfd plant:
Hydrogen—steam methane
reforming (SMR) ��
��
��
��
��
��
��
��
Application: Production of hydrogen from hydrocarbon feedstocks such
as: natural gas, LPG, butane, naphtha, reﬁnery offgases, etc., using the
Haldor Topsøe radiant-wall Steam Methane Reformer (SMR). Plant ca-
pacities range from 5,000 Nm3/h to more than 200,000 Nm3/h hydro-
gen (200+ MMscfd H2) and hydrogen purity of up to 99.999+%.

Description: The Haldor Topsøe SMR-based hydrogen plant is tailor-made ��

to suit the customer’s needs with respect to economics, feedstock flex- ��
ibility and steam export. In a typical Topsøe SMR-based hydrogen plant,
a mix of hydrocarbon feedstocks or a single feedstock stream is first de-
sulfurized. Subsequently, process steam is added, and the mixture is fed ��
to a prereformer. Further reforming is carried out in the Haldor Topsøe ��
radiant wall SMR. The process gas is reacted in a medium-temperature ��
CO shift reactor and purified by pressure swing absorption (PSA) to ob-
tain product-grade hydrogen. PSA offgases are used as fuel in the SMR.
Excess heat in the plant is efficiently used for process heating and steam
generation.
The Haldor Topsøe radiant wall SMR operates at high outlet temper-
References: Rostrup-Nielsen, J. R. and T. Rostrup-Nielsen, “Large scale
atures up to 950°C (1,740°F). The Topsøe reforming catalysts allow op-
hydrogen production,” CatTech, Vol. 6, no. 4, 2002.
eration at low steam-to-carbon ratio. Advanced Steam Reforming uses
Dybkjær, I., and S. W. Madsen, “Advanced reforming technologies
both high outlet temperature and low steam-to-carbon ratio, which are
for hydrogen production,” Hydrocarbon Engineering, December/Janu-
necessary for high-energy efficiency and low hydrogen production cost.
ary 1997/1998.
The Advanced Steam Reforming design is in operation in many indus-
Gøl, J.N., and I. Dybkjær, “Options for hydrogen production,” HTI
trial plants throughout the world.
Quarterly: Summer 1995.
Economics: The Advanced Steam Reforming conditions described can Licensor: Haldor Topsøe A/S.
achieve a net energy efficiency as low as 2.96 Gcal /1,000 Nm3 hydrogen
using natural gas feed (315 MM Btu/scf H2).

Installations: More than 100 units.

Hydroprocessing, residue
Application: Produces maximum distillates and low-sulfur fuel oil, or ��
low-sulfur LR-CCU feedstock, with very tight sulfur, vanadium and CCR
specifications, using moving bed “bunker” and fixed-bed technologies.
Bunker units are available as a retrofit option to existing fixed-bed resi-
��
due HDS units. ��

Description: At limited feed metal contents, the process typically uses

all fixed-bed reactors. With increasing feed metal content, one or more ��
moving-bed “bunker” reactors are added up-front of the fixed-bed re-
actors to ensure a fixed-bed catalyst life of at least one year. A steady
��
state is developed by continuous catalyst addition and withdrawal: the ��
catalyst aging is fully compensated by catalyst replacement, at typically
0.5 to 2 vol% of inventory per day. ��
��
An all bunker option, which eliminates the need for catalyst change-out, ��
��
is also available. A hydrocracking reactor, which converts the synthetic vacu-
um gasoil into distillates, can be efficiently integrated into the unit. A wide
range of residue feeds, like atmospheric or vacuum residues and deasphalted
oils, can be processed using Shell residue hydroprocessing technologies.
unit. Investment costs for a typical new single string 5,000 tpsd SR-Hy-
Operating conditions: con unit will range from 200 to 300 MM US$; the higher figure includes
Reactor pressures: 100 – 200 bar 1,450 – 3,000 psi an integrated hydrocracker.
Reactor temperatures: 370 – 420°C 700 – 790°F
Installation: There is one unit with both bunker reactors and fixed-bed
Yields: Typical yields for an SR HYCON unit on Kuwait feed: reactors, operating on short residue (vacuum residue) at 4,300 tpd or
Feedstock SR (95% 520C+) with integrated HCU 27 Mbpsd capacity, and two all-fixed bed units of 7,700 and 7,000 tpd
Yields: [%wof] [%wof] (48 and 44 Mbpsd resp.), the latter one in one single string. Commercial
Gases C1 – C4 3 5 experiences range from low-sulfur atmospheric residues to high-metal,
Naphtha C5 – 165°C 4 18 high-sulfur vacuum residues with over 300-ppmw metals.
Kero + gasoil 165 – 370°C 20 43
Reference: Scheffer, B., et al, “The Shell Residue Hydroconversion Pro-
VGO 370 – 580°C 41 4
cess: Development and achievements,” The European Refining Technol-
Residue 580°C+ 29 29
ogy Conference, London, November 1997.
H2 cons. 2 3
Licensor: Shell Global Solutions International B.V.
Economics: Investment costs for the various options depend strongly on
feed properties and process objectives of the residue hydroprocessing
Hydroprocessing, ULSD ��
��
Application: A versatile family of ExxonMobil Research and Engineering
Co. (EMRE) process technologies and catalysts are used to meet all cur- ��
��
��
rent and possible future premium diesel requirements.
��
ULSD HDS—Ultra-deep hydrodesulfurization process to produce distil-
��
late products with sulfur levels below 10 wppm. ��
HDHC—Heavy-distillate mild-hydrocracking process for the reduction of ��
��
T90 and T95 boiling points, and high-level density reduction. ��
��
MAXSAT—High-activity aromatics saturation process for the selective ��
reduction of polyaromatics under low pressure and tempera- ��
��
ture conditions.
��
CPI—Diesel cloud point improvement by selective normal paraffin hydro-
cracking (MDDW) or by paraffin isomerization dewaxing (MIDW). ��
��
Description: EMRE units combine the technologies listed above in ��
low-cost integrated designs to achieve the necessary product uplift at ��
minimum investment and operating cost. For ultra-low-sulfur-diesel
hydrodesulfurization (ULSD HDS), a single-stage single-reactor process
can be designed. A small cetane improvement, together with the reduc-
tion of polyaromatics to less than 11 wt.% or as low as 5 wt.%, can closely integrated with ULSD HDS and other functions to achieve the full
be economically achieved with proper specification of catalyst, hydrogen upgrading requirements.
partial pressure, space velocity and the installation of high-performance The EMRE ULSD technologies are equally amenable to revamp or
Spider Vortex internals. grassroots applications. EMRE has an alliance with Kellogg Brown &
The addition of heavy-diesel hydrocracking (HDHC) function to the Root (KBR) to provide these technologies to refiners.
HDS reactor can achieve T95 boiling point reduction together with
Economics:
higher levels of density and aromatics reduction and greater cetane
Investment: (Basis: 20,000–35,000 bpsd, 1st quarter
improvement.
2004 US Gulf Coast)
When feedstock aromatics are very high, or very low aromatics in
New unit, $/bpsd 1,200–2,000
the product are desired, a second-stage aromatics saturation (MAXSAT)
system is specified to avoid very high design pressures required for a Installation: Nineteen distillate upgrading units have applied the EMRE
single-step base-metal hydrotreating catalyst system. When the distil- ULSD technologies. Twelve of these applications are revamps.
late product must also meet stringent fluidity specifications, EMRE can
offer either paraffin isomerization dewaxing (MIDW) or selective normal Licensor: ExxonMobil Research and Engineering Co.
paraffin cracking-based dewaxing technologies (MDDW). These can be
Hydrotreating
Application: Hydroprocessing of middle distillates, including cracked
materials (coker/visbreaker gas oils and LCO), using SynTechnology ��
��
maximizes distillate yield while producing ultra-low-sulfur diesel (ULSD) ��
with improved cetane and API gain, reduced aromatics, T95 reduction ��
and cold-flow improvement through selective ring opening, saturation ��

and/or isomerization. Various process conﬁgurations are available for

revamps and new unit design to stage investments to meet changing
diesel specifications.
��
Products: Maximum yield of improved quality distillate while minimizing
fuel gas and naphtha. Diesel properties include less than 10-ppm sulfur,
with aromatics content (total and/or PNA), cetane, density and T95 de-
pendent on product objectives and feedstock. ��
��
Description: SynTechnology includes SynHDS for ultra-deep desulfuri-
zation and SynShift / SynSat for cetane improvement, aromatics satura-
tion and density / T95 reduction. SynFlow for cold flow improvement
can be added as required. The process combines ABB Lummus Glob-
al’s cocurrent and/or patented countercurrent reactor technology with
special SynCat catalysts from Criterion Catalyst Co. LP. It incorporates grading of a feed blend containing 72% LCO and LCGO gave these
design and operations experience from Shell Global Solutions to maxi- performance figures:
mize reactor performance by using advanced reactor internals.
A single-stage or integrated two-stage reactor system provides Feed blend Product
various process configuration options and revamp opportunities. In a Gravity, °API 25 33.1
two-stage reactor system, the feed, makeup and recycle gas are heated Sulfur, wt% (wppm) 1.52 (2)
and fed to a first-stage cocurrent reactor. Effluent from the first stage Nitrogen, wppm 631 <1
is stripped to remove impurities and light ends before being sent to the Aromatics, vol% 64.7 34.3
second-stage countercurrent reactor. When a countercurrent reactor Cetane index 34.2 43.7
is used, fresh makeup hydrogen can be introduced at the bottom of Liquid yield on feed, vol% 107.5
the catalyst bed to achieve optimum reaction conditions.
Economics: SynTechnology encompasses a family of low-to-moder-
Operating conditions: Typical operating conditions range from 500 – ate pressure processes. Investment cost will be greatly dependent on
1,000 psig and 600°F – 750°F. Feedstocks range from straight-run feed quality and hydroprocessing objectives. For a 30,000 to 35,000-
gas oils to feed blends containing up to 70% cracked feedstocks that
have been commercially processed. For example, the SynShift up- Continued 
Hydrotreating, continued
bpsd unit, the typical ISBL investment cost in US$/bpsd (2006 US Gulf
Coast) are:
Revamp existing unit 600 –1,300
New unit for deep HDS 1,500 –1,700
New unit for cetane improvement and HDA 2,100 –2,500

Installation: SynTechnology has been selected for more than 30 units,

with half of the projects being revamps. Twenty units are in operation.

Licensor: ABB Lummus Global, on behalf of the SynAlliance, which in-

cludes Criterion Catalyst and Technologies Co., and Shell Global Solu-
tions.
Hydrotreating
Application: Reduction of the sulfur, nitrogen and metals content of
naphthas, kerosines, diesel or gas oil streams. ��
��
Products: Low-sulfur products for sale or additional processing. ��

Description: Single or multibed catalytic treatment of hydrocarbon liq-

uids in the presence of hydrogen converts organic sulfur to hydrogen ��
��
sulﬁde and organic nitrogen to ammonia. Naphtha treating normally oc- ��
��
curs in the vapor phase, and heavier oils usually operate in mixed-phase. ��
��
Multiple beds may be placed in a single reactor shell for purposes of
redistribution and/or interbed quenching for heat removal. Hydrogen-
rich gas is usually recycled to the reactor(s) to maintain adequate hydro- ��
��
gen-to-feed ratio. Depending on the sulfur level in the feed, H2S may be ��
scrubbed from the recycle gas. Product stripping is done with either a ��
��
reboiler or with steam. Catalysts are cobalt-molybdenum, nickel-molyb-
denum, nickel-tungsten or a combination of the three.

Operating conditions: 550°F to 750°F and 400 psig to 1,500 psig reac-
tor conditions.

Yields: Depend on feed characteristics and product speciﬁcations. Re-

covery of desired product typically exceeds 98.5 wt% and usually ex-
ceeds 99%.

Economics:
Utilities, (per bbl feed) Naphtha Diesel
Fuel, 103 Btu release 48 59.5
Electricity, kWh 0.65 1.60
Water, cooling (20°F rise), gal 35 42

Licensor: CB&I Howe-Baker.

Hydrotreating
��
Application: The CDHydro and CDHDS processes are used to selectively
desulfurize FCC gasoline with minimum octane loss.
��
Products: Ultra-low-sulfur FCC gasoline with maximum retention of ��
oleﬁns and octane. ��

Description: The light, mid and heavy cat naphthas (LCN, MCN, HCN)
��
are treated separately, under optimal conditions for each. The full- ��
range FCC gasoline sulfur reduction begins with fractionation of the
light naphtha overhead in a CDHydro column. Mercaptan sulfur re-
��
acts quantitatively with excess diolefins to produce heavier sulfur com-
pounds, and the remaining diolefins are partially saturated to olefins ��
by reaction with hydrogen. Bottoms from the CDHydro column, con-
taining the reacted mercaptans, are fed to the CDHDS column where
the MCN and HCN are catalytically desulfurized in two separate zones.
HDS conditions are optimized for each fraction to achieve the desired
or to replace catalyst. Typical fixed-bed processes will require a mid FCC
sulfur reduction with minimal olefin saturation. Olefins are concentrat-
shutdown to regenerate/replace catalyst, requiring higher capital cost
ed at the top of the column, where conditions are mild, while sulfur
for feed, storage, pumping and additional feed capacity.
is concentrated at the bottom where the conditions result in very high
levels of HDS. Economics: The estimated ISBL capital cost for a 50,000-bpd CDHydro/
No cracking reactions occur at the mild conditions, so that yield CDHDS unit with 95% desulfurization is $40 million (2005 US Gulf Coast).
losses are easily minimized with vent-gas recovery. The three product Direct operating costs—including utilities, catalyst, hydrogen and octane
streams are stabilized together or separately, as desired, resulting in replacement—are estimated at $0.04/gal of full-range FCC gasoline.
product streams appropriate for their subsequent use. The two columns
are heat integrated to minimize energy requirements. Typical reformer Installation: Twenty-one CDHydro/CDHDS units are in operation treating
hydrogen is used in both columns without makeup compression. The FCC gasoline and 12 more units are currently in engineering/construc-
sulfur reduction achieved will allow the blending of gasoline that meets tion. Total licensed capacity exceeds 1.3 million bpd.
current and future regulations.
Catalytic distillation essentially eliminates catalyst fouling because Licensor: CDTECH.
the fractionation removes heavy-coke precursors from the catalyst zone
before coke can form and foul the catalyst pores. Thus, catalyst life in
catalytic distillation is increased significantly beyond typical fixed-bed
life. The CDHydro/CDHDS units can operate throughout an FCC turn-
around cycle up to six years without requiring a shutdown to regenerate
Hydrotreating ��
Application: Topsøe hydrotreating technology has a wide range of ap- ��
��
plications, including the purification of naphtha, distillates and residue,
as well as the deep desulfurization and color improvement of diesel fuel ��
��
and pretreatment of FCC and hydrocracker feedstocks.

Products: Ultra-low-sulfur diesel fuel, and clean feedstocks for FCC and ��
hydrocracker units.
��

Description: Topsøe’s hydrotreating process design incorporates our in- ��
dustrially proven high-activity TK catalysts with optimized graded-bed
loading and high-performance, patented reactor internals. The combi- ��
��
nation of these features and custom design of grassroots and revamp ��
��
hydrotreating units result in process solutions that meet the refiner’s ��
objectives in the most economic way. ��
In the Topsøe hydrotreater, feed is mixed with hydrogen, heated
and partially evaporated in a feed/effluent exchanger before it enters
the reactor. In the reactor, Topsøe’s high-efficiency internals have a
low sensitivity to unlevelness and are designed to ensure the most Operating conditions: Typical operating pressures range from 20 to 80
effective mixing of liquid and vapor streams and the maximum utili- barg (300 to 1,200 psig), and typical operating temperatures range from
zation of the catalyst volume. These internals are effective at a high 320°C to 400°C (600°F to 750°F).
range of liquid loadings, thereby enabling high turndown ratios. Top-
References: Cooper, B. H. and K. G. Knudsen, “Production of ULSD: Cat-
søe’s graded-bed technology and the use of shape-optimized inert
alyst, kinetics and reactor design,” World Petroleum Congress, 2002.
topping and catalysts minimize the build-up of pressure drop, there-
Patel, R. and K. G. Knudsen, “How are refiners meeting the ul-
by enabling longer catalyst cycle length. The hydrotreating catalysts
tra-low-sulfur diesel challenge,” NPRA Annual Meeting, San Antonio,
themselves are of the Topsøe TK series, and have proven their high
March 2003.
activities and outstanding performance in numerous operating units
Topsøe, H., K. Knudsen, L. Skyum and B. Cooper, “ULSD with BRIM
throughout the world. The reactor effluent is cooled in the feed-ef-
catalyst technology,” NPRA Annual Meeting, San Francisco, March 2005.
fluent exchanger, and the gas and liquid are separated. The hydro-
gen gas is sent to an amine wash for removal of hydrogen sulfide Installation: More than 60 Topsøe hydrotreating units for the various ap-
and is then recycled to the reactor. Cold hydrogen recycle is used as plications above are in operation or in the design phase.
quench gas between the catalyst beds, if required. The liquid product
is steam stripped in a product stripper column to remove hydrogen Licensor: Haldor Topsøe A/S.
sulfide, dissolved gases and light ends.
Hydrotreating
Application: The IsoTherming process provides refiners an economical ��
means to produce ultra-low-sulfur diesel (ULSD), low-sulfur and low-
nitrogen FCC feedstocks, and other very low-sulfur hydrocarbon prod- � �
��
ucts. In addition, IsoTherming can provide a cost-effective approach to
wax and petrolatum hydrogenation to produce food-grade or pharma-
ceutical-grade oil and wax products, and lubestock hydroprocessing for
sulfur reduction and VI improvement. � � �

Products: ULSD, low-sulfur FCC feed, low-sulfur gasoline-kerosine type

products. High-quality lube-oil stock and food- and pharmaceutical- �
grade oil and wax. ��

Description: This process uses a novel approach to introduce hydrogen ��

into the reactor; it enables much higher space velocities than conven-
tional hydrotreating reactors. The IsoTherming process removes the
hydrogen mass transfer limitation and operates in a kinetically limited
mode since hydrogen is delivered to the reactor in the liquid phase as
soluble hydrogen. Operating conditions: Typical diesel IsoTherming conditions are:
The technology can be installed as a simple pre-treat unit ahead Diesel feed IsoTherming Treated product
of an existing hydrotreater reactor or a new stand-alone process unit. pre-treat reactor from existing
Fresh feed, after heat exchange, is combined with hydrogen in Reactor
conventional
One mixer (1). The feed liquid with soluble hydrogen is fed to IsoTherm-
ing Reactor One (2) where partial desulfurization occurs. The stream is
reactor
LCO, vol% 40
combined with additional hydrogen in Reactor Two Mixer (3), and fed to
SR, vol% 60
IsoTherming Reactor Two (4) where further desulfurization takes place.
Sulfur, ppm 7,500 900 5
Treated oil is recycled (5) back to the inlet of Reactor One. This re-
Nitrogen, ppm 450 50 0
cycle stream delivers more hydrogen to the reactors and also acts as a
H2 consumption, scf/bbl 300 150
heat sink; thus, a nearly isothermal reactor operation is achieved.
LHSV, Hr –1* 5 2.5
The treated oil from IsoTherming Reactor Two (4) may then be fed
Reactor T 30 30
to additional IsoTherming reactors and / or to a trickle hydrotreating re-
Reactor pressure, psig 1,110 900
actor (6) in the polishing mode to produce an ultra-low-sulfur product.
*Based on fresh feedrate without recycle

Continued 
Hydrotreating, continued
Economics: Revamp investment (basis 15,000 –20,000 bpsd, 1Q 2004,
US Gulf Coast) $400/bpsd diesel

Installation: Four units have been licensed for ULSD; two units licensed
for gasoil mild hydrocracking.

Licensor: P. D. Licensing, LLC (Process Dynamics, Inc.).

Hydrotreating
Application: Hydrodesulfurization, hydrodenitrogenation and hydro- ��
genation of petroleum and chemical feedstocks using the Unionﬁning
��
and MQD Unionﬁning processes. ��

Products: Ultra-low-sulfur diesel fuel; feed for catalytic reforming, �

�
FCC pretreat; upgrading distillates (higher cetane, lower aromatics);
desulfurization, denitrogenation and demetallization of vacuum and at-
mospheric gas oils, coker gas oils and chemical feedstocks. ��
��
Description: Feed and hydrogen-rich gas are mixed, heated and contact- ��
ed with regenerable catalyst (1). Reactor efﬂuent is cooled and separated
(2). Hydrogen-rich gas is recycled or used elsewhere. Liquid is stripped
(3) to remove light components and remaining hydrogen sulﬁde, or frac-
tionated for splitting into multiple products.

Operating conditions: Operating conditions depend on feedstock and

desired level of impurities removal. Pressures range from 500 to 2,000 Distillate, vol% 97.2 97.6 98.0 99.0
psi. Temperatures and space velocities are determined by process objec- Gravity, °API 24.0 26.9 27.8 35.2
tives. Boiling range, °F 400+ 325/660 300+ 300
Sulfur, wt% 0.025 0.001 0.002 0.001
Yields: H2 consump., scf/bbl 700 350 620 300
Purpose FCC feed Desulf. Desulf. Desulf.
Feed, source VGO + Coker AGO VGO DSL Economics:
Gravity, °API 17.0 25.7 24.3 32.9 Investment, $ per bpsd 1,200–2,000
Boiling range, °F 400/1,000 310/660 540/1,085 380/700 Utilities, typical per bbl feed:
Sulfur, wt% 1.37 1.40 3 1.1 Fuel, 103 Btu 40–100
Nitrogen, ppmw 6,050 400 1,670 102 Electricity, kWh 0.5–1.5
Bromine number — 26 — —
Installation: Several hundred units installed.
Naphtha, vol% 4.8 4.2 3.9 1.6
Gravity, °API 45.0 50.0 54.0 51 Licensor: UOP LLC.
Boiling range, °F 180/400 C4/325 C4/356 C5/300
Sulfur, ppmw 50 <2 <2 <1
Nitrogen, ppmw 30 <1 <2 <0.5
Hydrotreating
Application: RCD Unionfining process reduces the sulfur, nitrogen, Con- ��
radson carbon, asphaltene and organometallic contents of heavier resi- ��
due-derived feedstocks to allow them to be used as either specification
��
fuel oils or as feedstocks for downstream processing units such as hydro- ��
crackers, fluidized catalytic crackers, resid catalytic crackers and cokers. ��
��
Feed: Feedstocks range from solvent-derived materials to atmospheric ��
and vacuum residues. ��
��
��
Description: The process uses a fixed-bed catalytic system that operates at ��
moderate temperatures and moderate to high hydrogen partial pressures. ��
Typically, moderate levels of hydrogen are consumed with minimal pro-
��
duction of light gaseous and liquid products. However, adjustments can ��
��
be made to the unit’s operating conditions, flowscheme configuration or ��
catalysts to increase conversion to distillate and lighter products. ��
Fresh feed is combined with makeup hydrogen and recycled gas,
and then heated by exchange and fired heaters before entering the
unit’s reactor section. Simple downflow reactors incorporating a graded
bed catalyst system designed to accomplish the desired reactions while
minimizing side reactions and pressure drop buildup are used. Reactor
effluent flows to a series of separators to recover recycle gas and liquid Installation: Twenty-six licensed units with a combined licensed capacity
products. The hydrogen-rich recycle gas is scrubbed to remove H2S and of approximately 900,000 bpsd. Commercial applications have included
recycled to the reactors while finished products are recovered in the processing of atmospheric and vacuum residues and solvent-derived
fractionation section. Fractionation facilities may be designed to simply feedstocks.
recover a full-boiling range product or to recover individual fractions of Licensor: UOP LLC.
the hydrotreated product.
Economics:
Investment (basis: 15,000 – 20,000 bpsd, 2Q 2002, US Gulf Coast)
$ per bpsd 2,000 – 3,500
Utilities, typical per barrel of fresh feed (20,000 bpsd basis)
Fuel, MMBtu/hr 46
Electricity, kWh 5,100
Steam, HP, lb / hr 8,900
Steam, LP, lb / hr 1,500
Hydrotreating ��
��
Application: Hydrotreating of light and middle distillates and various gas
oils, including cracked feedstocks (coker naphtha, coker LGO and HGO, ��
visbreaker gas oil, and LCO) using the ISOTREATING Process for deep
desulfurization, denitrification and aromatics saturation and to produce ��
�
low-sulfur naphtha, jet fuel, ultra-low sulfur diesel (ULSD), or improved- ��
��
quality FCC feed. � ��
��
�
� �
Description: Feedstock is mixed with hydrogen-rich treat gas, heated ��
��
and reacted over high-activity hydrogenation catalyst (1). Several CoMo � �
and NiMo catalysts are available for use in the ISOTREATING Process. One �
��
or multiple beds of catalyst(s), together with Chevron Lummus Global’s �
��
advanced high-efficiency reactor internals for reactant distribution and
� ��
interbed quenching, are used. ��
Reactor effluent is cooled and flashed (2) producing hydrogen-rich ��
recycle gas, which, after H2S removal by amine (3), is partially used as
quench gas while the rest is combined with makeup hydrogen gas to
form the required treat gas. An intermediate pressure level flash (4) can to produce ULSD (<10 wppm sulfur). Chemical-hydrogen consumption
be used to recover some additional hydrogen-rich gas from the liquid ranges from 450 –900+ scf/bbl feed.
effluent prior to the flashed liquids being stripped or fractionated (5) to
remove light ends, H2S and naphtha-boiling range material, and/or to Economics: Investment will vary depending on feedstock characteristics
fractionate the higher boiling range materials into separate products. and product requirements. For a 40,000–45,000-bpsd unit for ULSD,
the ISBL investment cost (US Gulf Coast 2006) is $700 –1,000/ bpsd for
Operating conditions: Typical reactor operating conditions can range a revamped unit and $1,700 –1,900/ bpsd for a new unit.
from 600 –2,300 psig and 500 –780°F, 350 –2,000 psia hydrogen partial
pressure, and 0.6-3 hr –1 LHSV, all depending on feedstock(s) and prod- Installation: Currently, there are more than 60 units operating based on
uct quality objective(s). ISOTREATING technology and an additional 10 units in various stages of
engineering.
Yields: Depends on feedstock(s) characteristics and product require-
ments. Desired product recovery is maximized based on required flash Licensor: Chevron Lummus Global LLC.
point and/or specific fractionation specification. Reactor liquid product
(350°F plus TBP material) is maximized through efficient hydrogenation
with minimum lighter liquid product and gas production. Reactor liq-
uid product (350°F plus) yield can vary between 98 vol% from straight-
run gas oil feed to >104 vol% from predominantly cracked feedstock
Hydrotreating, diesel
Application: Produce ultra-low-sulfur diesel (ULSD) and high-quality die-
sel fuel (low aromatics, high cetane) via Prime-D toolbox of proven state- � ��
of-the-art technology, catalysts and services. �
��
��
��
Description: In the basic process, as shown in the diagram, feed and
�
hydrogen are heated in the feed-reactor effluent exchanger (1) and fur-
nace (2) and enter the reaction section (3), with possible added volume �
��
for revamp cases. The reaction effluent is cooled by the exchanger (1) � ��
�
and air cooler (4) and separated in the separator (5). The hydrogen-
rich gas phase is treated in an existing or new amine absorber for H2S �

removal (6) and recycled to the reactor. The liquid phase is sent to the ��
��
stripper (7) where small amounts of gas and naphtha are removed and ��
high-quality product diesel is recovered. ��
��
Whether the need is for a new unit or for maximum reuse of existing
diesel HDS units, the Prime-D hydrotreating toolbox of solutions meets
the challenge. Process objectives ranging from low-sulfur, ultra-low-sul-
fur, low-aromatics, and/or high cetane number are met with minimum
cost by:
• Selection of the proper catalyst from the HR 500 Series, based on and technical service feedback to ensure the right application of the
the feed analysis and processing objectives. HR 500 catalysts cover the right technology for new and revamp projects.
range of ULSD requirements with highly active and stable catalysts. HR Whatever the diesel quality goals—ULSD, high cetane or low aro-
526 CoMo exhibits high desulfurization rates at low to medium pres- matics—Prime-D’s Hydrotreating Toolbox approach will attain your goals
sures; HR 538/HR 548 NiMo have higher hydrogenation activities at in a cost-effective manner.
higher pressures. Installation: Over 150 middle distillate hydrotreaters have been li-
• Use of proven, efficient reactor internals, EquiFlow, that allow censed or revamped. They include 56 low- and ultra-low-sulfur diesel
near-perfect gas and liquid distribution and outstanding radial tempera- units (<50 ppm), as well as a number of cetane boosting units. Most
ture profiles. of those units are equipped with Equiflow internals.
• Loading catalyst in the reactor(s) with the Catapac dense loading
technique for up to 20% more reactor capacity. Over 10,000 tons of References: “Getting Total Performance with Hydrotreating,” Petroleum
catalyst have been loaded quickly, easily and safely in recent years using Technology Quarterly, Spring 2002.
the Catapac technique. “Premium Performance Hydrotreating with Axens HR 400 Series
• Application of Advanced Process Control for dependable opera- Hydrotreating Catalysts,” NPRA Annual Meeting, March 2002, San
tion and longer catalyst life. Antonio.
• Sound engineering design based on years of R&D, process design Continued 
Hydrotreating, diesel, continued
“The Hydrotreating Toolbox Approach,” Hart’s European Fuel News,
May 29, 2002.
“Squeezing the most from hydrotreaters,” Hydrocarbon Asia, April/
May 2004.

Licensor: Axens.
Hydrotreating/desulfurization
Application: The SelectFining process is a gasoline desulfurization tech- ��
nology developed to produce ultra-low-sulfur gasoline by removing ��
more than 99% of the sulfur present in oleﬁnic naphtha while: ��
• Minimizing octane loss ��

• Maximizing liquid yield

��
• Minimizing H2 consumption
• Eliminating recombination sulfur.
��
Description: The SelectFining process can hydrotreat full boiling-range
(FBR) olefinic naphtha or, when used in conjunction with a naphtha split-
ter, any fraction of FBR naphtha.
The configuration of a single-stage SelectFining unit processing
FBR olefinic naphtha (Fig. 1) is very similar to that of a conventional ��
��
hydrotreater. The operating conditions of the SelectFining process are
similar to those of conventional hydrotreating: it enables refiners to re-
use existing idle hydroprocessing equipment.
Since FBR olefinic naphtha can contain highly reactive di-olefins
(which may polymerize and foul equipment and catalyst beds), the Se-
lectFining unit may include a separate reactor for di-olefin stabilization. port (with optimized acidity) and non-noble metal promoters to achieve
Incoming naphtha is mixed with a small stream of heated hydrogen-rich the optimal combination of desulfurization, olefin retention and operat-
recycle gas and directed to this reactor. The “stabilized” naphtha is then ing stability.
heated to final reaction conditions and processed in the unit’s main reac- In addition to processing FBR naphtha, the SelectFining technology
tor over SelectFining catalyst. can also be used in an integrated gasoline upgrading configuration that
Effluent from the main reactor is washed, cooled and separated into includes naphtha splitting, Merox extraction technology for mercaptan
liquid and gaseous fractions. Recovered gases are scrubbed (for H2S re- removal and ISAL hydroconversion technology for octane recovery.
moval) and recycled to the unit’s reactor section, while recovered liquids Economics: A SelectFining unit can preserve olefins while desulfurizing
are debutanized (for Rvp control) and sent to gasoline blending. FBR naphtha from a fluid catalytic cracking unit. When producing a 50-
Process chemistry: While the principal reactions that occur in a wppm sulfur product (~98% HDS), the additional olefin retention pro-
hydrotreater involve conversion of sulfur and nitrogen components, vided by the single-stage SelectFining unit corresponds to a 2.5 (R+M)/2
conventional hydrotreaters also promote other reactions, including octane advantage relative to conventional hydrotreating. This advantage
olefin saturation, reducing the feed’s octane. UOP’s S 200 SelectFining increases to 3.5 (R+M)/2 when a two-stage unit is used. Based upon an
catalyst was developed to effectively hydrotreat the olefinic naphtha octane value of $0.25 per octane-bbl, hydrogen cost of $3 per 1,000
while minimizing olefin saturation. It uses an amorphous alumina sup- Continued 
Hydrotreating/desulfurization, continued
SCFB and 20,000 BPSD naphtha throughput, the resulting savings in
processing costs can range from $4 to $6 million per year depending
upon the SelectFining process flowscheme applied.

Installation: UOP’s experience in hydroprocessing and gasoline

desulfurization is extensive with approximately 200 Unionﬁning units
and more than 240 Merox units (for naphtha service) in operation.

Licensor: UOP LLC.

Hydrotreating—aromatic saturation
Application: Hydroprocessing of middle distillates, including cracked ��
materials (coker/visbreaker gas oils and LCO), using SynTechnology ��
maximizes distillate yield while producing ultra-low-sulfur diesel (ULSD) ��
��
with improved cetane and API gain, reduced aromatics, T95 reduction ��
and cold-flow improvement through selective ring opening, saturation
and/or isomerization. Various process configurations are available for
revamps and new unit design to stage investments to meet changing
diesel specifications.
��
Products: Maximum yield of improved quality distillate while minimizing
fuel gas and naphtha. Diesel properties include less than 10-ppm sulfur,
with aromatics content (total and/or PNA), cetane, density and T95 de-
pendent on product objectives and feedstock. ��
��
Description: SynTechnology includes SynHDS for ultra-deep desulfuri-
zation and SynShift/SynSat for cetane improvement, aromatics satura-
tion and density/T95 reduction. SynFlow for cold flow improvement can
be added as required. The process combines ABB Lummus Global’s co-
current and/or patented countercurrent reactor technology with special been commercially processed. For example, the SynShift upgrading of
SynCat catalysts from Criterion Catalyst Co. LP. It incorporates design a feed blend containing 72% LCO and LCGO gave these performance
and operations experience from Shell Global Solutions to maximize reac- figures:
tor performance by using advanced reactor internals.
A single-stage or integrated two-stage reactor system provides Feed blend Product
various process configuration options and revamp opportunities. In a Gravity, °API 25 33.1
two-stage reactor system, the feed, makeup and recycle gas are heated Sulfur, wt% (wppm) 1.52 (2)
and fed to a first-stage cocurrent reactor. Effluent from the first stage Nitrogen, wppm 631 <1
is stripped to remove impurities and light ends before being sent to the Aromatics, vol% 64.7 34.3
second-stage countercurrent reactor. When a countercurrent reactor is Cetane index 34.2 43.7
used, fresh makeup hydrogen can be introduced at the bottom of the Liquid yield on feed, vol% 107.5
catalyst bed to achieve optimum reaction conditions. Economics: SynTechnology encompasses a family of low-to-moderate
Operating conditions: Typical operating conditions range from 500 – pressure processes. Investment cost will be greatly dependent on feed
1,000 psig and 600°F – 750°F. Feedstocks range from straight-run gas quality and hydroprocessing objectives. For a 30,000 to 35,000-bpsd unit,
oils to feed blends containing up to 70% cracked feedstocks that have the typical ISBL investment cost in US$/bpsd (2006 US Gulf Coast) are:
Continued 
Hydrotreating—aromatic saturation, continued
Revamp existing unit 600 –1,300
New unit for deep HDS 1,500 –1,700
New unit for cetane improvement and HDA 2,100 –2,500

Installation: SynTechnology has been selected for more than 30 units,

with half of the projects being revamps. Twenty units are in operation.

Licensor: ABB Lummus Global, on behalf of the SynAlliance, which in-

cludes Criterion Catalyst and Technologies Co., and Shell Global Solu-
tions.
Hydrotreating—lube and wax
Application: The IsoTherming process provides reﬁners with a cost- ��
��
effective approach to lube and wax hydrotreating to produce high-qual-
ity lube-base stocks and food-grade waxes.

Products: High-quality lube oils, food- or pharmaceutical-grade oil and

wax products.

Description: This process uses a novel approach to introduce hydro- ��

gen into the reactor; it enables much higher LHSV than conventional ��
hydrotreating reactors. The IsoTherming process removes the hydro-
gen mass transfer limitation and operates in a kinetically limited mode
since hydrogen is delivered to the reactor in the liquid phase as soluble
hydrogen.
The technology can be installed as a simple pre-treat unit ahead of ��
an existing hydrotreater reactor or a new stand-alone process unit. Fresh
feed, after heat exchange, is combined with hydrogen in Reactor One
mixer (1). The liquid feed with soluble hydrogen is fed to IsoTherming
Reactor One (2) where partial desulfurization, denitroﬁcation and satu-
ration occurs.
The stream is combined with additional hydrogen in Reactor
Two mixer (3), and fed to IsoTherming Reactor Two (4) where further
desulfurization, denitroﬁcation and saturation take place. Treated oil is
recycled (5) to the inlet of Reactor One. This recycle stream delivers more
hydrogen to the reactors and also acts as a heat sink; thus, a nearly iso-
thermal reactor operation is achieved.

Economics: Investment (basis 5,000 bpd)

Lube-base oil grassroots $7.5 million
Lube-base oil retroﬁt $3.1 million
Parafﬁn wax grassroots $6.0 million
Micro-wax grassroots $13.0 million
White oil $10.7 million

Licensor: P. D. Licensing, LLC (Process Dynamics, Inc.).

Hydrotreating—RDS/VRDS/UFR/OCR
��
Application: Hydrotreat atmospheric and vacuum residuum feedstocks
to reduce sulfur, metals, nitrogen, carbon residue and asphaltene con-
��
tents. The process converts residuum into lighter products while im- ��
��
proving the quality of unconverted bottoms for more economic down-
��
stream use. ��
��
Products: Residuum FCC feedstock, coker feedstock, SDA feedstock or ��
��
low-sulfur fuel oil. VGO product, if separated, is suitable for further up- ��
grading by FCC units or hydrocrackers for gasoline/mid-distillate manu- ��
��
facture. Mid-distillate material can be directly blended into low-sulfur
��
diesel or further hydrotreated into ultra-low-sulfur diesel (ULSD).
��
The process integrates well with residuum FCC units to minimize ��
catalyst consumption, improve yields and reduce sulfur content of FCC
products. RDS/VRDS also can be used to substantially improve the yields
of downstream cokers and SDA units.

Description: Oil feed and hydrogen are charged to the reactors in a suited to revamp existing RDS/VRDS units for additional throughput or
once-through operation. The catalyst combination can be varied signifi- heavier feedstock.
cantly according to feedstock properties to meet the required product
Installation: Over 26 RDS/VRDS units are in operation. Six units have ex-
qualities. Product separation is done by the hot separator, cold separator
tensive experience with VR feedstocks. Sixteen units prepare feedstock
and fractionator. Recycle hydrogen passes through an H2S absorber.
for RFCC units. Four OCR units and two UFR unit are in operation, with
A wide range of AR, VR and DAO feedstocks can be processed. Ex-
another six in engineering. Total current operating capacity is about 1.1
isting units have processed feedstocks with viscosities as high as 6,000
million bpsd
cSt at 100°C and feed-metals contents of 500 ppm.
Onstream Catalyst Replacement (OCR) reactor technology has been References: Reynolds, “Resid Hydroprocessing With Chevron Technol-
commercialized to improve catalyst utilization and increase run length ogy,” JPI, Tokyo, Japan, Fall 1998.
with high-metals, heavy feedstocks. This technology allows spent cata- Reynolds and Brossard, “RDS/VRDS Hydrotreating Broadens Appli-
lyst to be removed from one or more reactors and replaced with fresh cation of RFCC,” HTI Quarterly, Winter 1995/96.
while the reactors continue to operate normally. The novel use of up- Reynolds, et al., “VRDS for conversion to middle distillate,” NPRA
flow reactors in OCR provides greatly increased tolerance of feed solids Annual Meetng, March 1998, Paper AM-98-23.
while maintaining low-pressure drop.
A related technology called UFR (upflow reactor) uses a multibed Licensor: Chevron Lummus Global LLC.
upflow reactor for minimum pressure drop in cases where onstream
catalyst replacement is not necessary. OCR and UFR are particularly well
Hydrotreating—resid
��
Application: Upgrade or convert atmospheric and vacuum residues us-
ing the Hyvahl fixed-bed process.

Products: Low-sulfur fuels (0.3% to 1.0% sulfur) and RFCC feeds (re-
moval of metals, sulfur and nitrogen, reduction of carbon residue). Thirty ��
percent to 50% conversion of the 565°C+ fraction into distillates. ��
��
Description: Residue feed and hydrogen, heated in a feed/efﬂuent ex-
changer and furnace, enter a reactor section—typically comprising of a
guard-reactor section, main HDM and HDS reactors.
The guard reactors are onstream at the same time in series, and they ��

protect downstream reactors by removing or converting sediment, met- ��

als and asphaltenes. For heavy feeds, they are permutable in operation
(PRS technology) and allow catalyst reloading during the run. Permuta-
tion frequency is adjusted according to feed-metals content and process
objectives. Regular catalyst changeout allows a high and constant pro- Utilities, per bbl feed:
tection of downstream reactors. Fuel, equiv. fuel oil, kg 0.3
Following the guard reactors, the HDM section carries out the re- Power, kWhr 10
maining demetallization and conversion functions. With most of the Steam production, MP, kg 25
contaminants removed, the residue is sent to the HDS section where the Steam consumption, HP, kg 10
sulfur level is reduced to the design speciﬁcation. Water, cooling, m3 1.1
The PRS technology associated with the high stability of the HDS
catalytic system leads to cycle runs exceeding a year even when process- Installation: In addition to three units in operation, three more were
ing VR-type feeds to produce ultra-low-sulfur fuel oil. licensed in 2005 / 06. Total installed capacity will reach 319,000 bpsd.
Two units will be operating on AR and VR feed, four on VR alone.
Yields: Typical HDS and HDM rates are above 90%. Net production of
12% to 25% of diesel + naphtha. References: Plain, C., D. Guillaume and E. Benaezi, “Better margins with
cheaper crudes,” ERTC 2005 Show Daily.
Economics: “Option for Resid Conversion,” BBTC, Oct. 8 – 9, 2002, Istanbul.
Investments (Basis: 40,000 bpsd, AR to VR feeds, “Maintaining on-spec products with residue hydroprocessing,”
2002 Gulf coast), US$/ bpsd 3,500–5,500 2000 NPRA Annual Meeting, March 26–28, 2000, San Antonio.

Licensor: Axens.
Hydrotreating—residue
Application: Upgrading or converting atmospheric and vacuum residues ��
using the Genoil GHU process.
Products: Removal of metals, maximize desulfurization (>90%), deni- ��
trogenation (>70%), reduction of carbon residue (>90%) with API in- ��
��
crease during the conversion process. Up to 90% conversion of 350°C+
fraction into distillates.
Description: Genoil has developed devices that enhance the mixing of
��
liquid hydrocarbons as well as highly efficient reactor internals. These
modifications have contributed to the high level of residue conversion,
desulfurization, denitrogenation and turnaround time.
Residue feed and hydrogen are heated in feed effluent exchanger
and furnace and enter the reactor section. The reactor section is typically
Investment: Based on 20,000 bpd, atmospheric or vacuum or VR feeds,
comprised of a guard-section hydrodemineralization (HDM) reactor, and
$3,000 – 5,000/bbl based of Gulf Coast rates.
sulfur-removal section—hydrodesulfurization (HDS).
The guard-reactor section protects the downstream HDS reactors by Installations: Genoil owns and operates a 10 BPD demonstration plant
removing metals and asphaltenes. Catalyst and operational objectives can at Two Hill, Alberta where we have conducted testing on several differ-
be adjusted according to feed metals content through different mechani- ent types of residue and crudes shipped to our facility by various com-
cal means to insure longer runtimes and constant protection for down- panies. We are currently working with several companies to get first
stream reactors. commercial installation of the GHU process.
After the guard section carries out the final removal of metals and
conversion, the HDS section removes the sulfur to design specifications. Reference: Asia Pacific Refining Conference Bangkok, Thailand, Sep-
With these factors in mind, Genoil has come up with a hydroconversion tember, 2005.
process that provides higher conversion, higher desulfurization and de- RPBC Moscow, Russia, April Conference 2006.
nitrogenation rates at lower pressure by using a simple, easy to operate Middle East Refining Conference, Doha, Qatar, May 2006.
process. America Oil and Gas Reporter, October 2005.
The upgrading process can be used as a field upgrader where heavy Energy Magazine, June 2005, Petroleum Technology Quarterly, Jan-
oil can be upgraded to WTI specification, pipeline specifications and pipe- uary 2006.
lined to refineries. Unconverted oil from the GHUunit can be sold as a Oil & Gas Product News, “Surge Global Announcement,” Jan./Feb.
stable, low-sulfur fuel oil or sent to another heavy-oil conversion unit for 2005.
further upgrading.
Licensor: Genoil Inc.
Yields: Net increase in production from 0 –10% naphtha, 1–20% kero-
sine and 21– 47% diesel.
Isomerization
Application: C5 /C6 paraffin-rich hydrocarbon streams are isomerized to
produce high RON and MON product suitable for addition to the gaso-
line pool. ��
��
Description: Several variations of the C5 / C6 isomerization process are ��
available. The choice can be a once-through reaction for an inexpensive-
but-limited octane boost, or, for substantial octane improvement and as ��

�
an alternate (in addition) to the conventional DIH recycle option, the Ip- ��
� �
sorb Isom scheme shown to recycle the normal paraffins for their com-
plete conversion. The Hexorb Isom configuration achieves a complete
normal paraffin conversion plus substantial conversion of low (75) oc- ��
tane methyl pentanes gives the maximum octane results. With the most
active isomerization catalyst (chlorinated alumina), particularly with the ��
Albemarle /Axens jointly developed ATIS2L catalyst, the isomerization
performance varies from 84 to 92: once-through isomerization -84,
isomerization with DIH recycle -88, Ipsorb -90, Hexorb-92.

Operating conditions: The Ipsorb Isom process uses a deisopentanizer

(1) to separate the isopentane from the reactor feed. A small amount of Power, kWh /h 310
hydrogen is also added to reactor (2) feed. The isomerization reaction Cooling water, m3/ h 100
proceeds at moderate temperature producing an equilibrium mixture of * Mid-2002, Gulf coast, excluding cost of noble metals
normal and isoparaffins. The catalyst has a long service life. The reactor
Installation: Of 35 licenses issued for C5 / C6 isomerization plants, 14
products are separated into isomerate product and normal paraffins in
units are operating including one Ipsorb unit.
the Ipsorb molecular sieve separation section (3) which features a novel
vapor phase PSA technique. This enables the product to consist entirely Reference: Axens /Albemarle, “Advanced solutions for paraffin
of branched isomers. isomerization,” NPRA Annual Meeting, March 2004, San Antonio.
“Paraffins isomerizatioin options,” Petroleum Technology Quarterly,
Economics: (Basis: Ipsorb “A” Isomerization unit with a 5,000-bpsd 70 Q2, 2005.
RONC feed needing a 20-point octane boost):
Investment*, million US$ 13.6 Licensor: Axens.
Utilities:
Steam, HP, tph 1.0
Steam, MP, tph 8.5
Steam, LP, tph 6.8
Isomerization
�
Application: Convert normal olefins to isoolefins. ��

Description:
C4 olefin skeletal isomerization (ISOMPLUS)
�
A zeolite-based catalyst especially developed for this process pro-
vides near equilibrium conversion of normal butenes to isobutylene at � �
high selectivity and long process cycle times. A simple process scheme
and moderate process conditions result in low capital and operating
costs. Hydrocarbon feed containing n-butenes, such as C4 raffinate, can ��
be processed without steam or other diluents, nor the addition of cata- ��
lyst activation agents to promote the reaction. Near-equilibrium con-
version levels up to 44% of the contained n-butenes are achieved at
greater than 90% selectivity to isobutylene. During the process cycle,
coke gradually builds up on the catalyst, reducing the isomerization ac-
tivity. At the end of the process cycle, the feed is switched to a fresh
catalyst bed, and the spent catalyst bed is regenerated by oxidizing the Total installed cost: Feedrate, Mbpd ISBL cost, $MM
coke with an air/nitrogen mixture. The butene isomerate is suitable for 10 8
making high purity isobutylene product. 15 11
30 20
C5 olefin skeletal isomerization (ISOMPLUS) Utility consumption: per barrel of feed (assuming an electric-motor-
A zeolite-based catalyst especially developed for this process pro- driven compressor) are:
vides near-equilibrium conversion of normal pentenes to isoamylene at Power, kWh 3.2
high selectivity and long process cycle times. Hydrocarbon feeds con- Fuel gas, MMBtu 0.44
taining n-pentenes, such as C5 raffinate, are processed in the skeletal Steam, MP, MMBtu 0.002
isomerization reactor without steam or other diluents, nor the addition Water, cooling, MMBtu 0.051
of catalyst activation agents to promote the reaction. Near-equilibrium Nitrogen, scf 57–250
conversion levels up to 72% of the contained normal pentenes are ob-
served at greater than 95% selectivity to isoamylenes. Installation: Two plants are in operation. Two licensed units are in vari-
ous stages of design.
Economics: The ISOMPLUS process offers the advantages of low capital
investment and operating costs coupled with a high yield of isobutylene Licensor: CDTECH and Lyondell Chemical Co.
or isoamylene. Also, the small quantity of heavy byproducts formed can
easily be blended into the gasoline pool. Capital costs (equipment, labor
and detailed engineering) for three different plant sizes are:
Isomerization
Application: Hydrisom is the ConocoPhillips selective diolefin hydroge-
nation process, with specific isomerization of butene-1 to butene-2 and ��
��
3-methyl-butene-1 to 2-methyl-butene-1 and 2-methyl-butene-2. The ��
Hydrisom Process uses a liquid-phase reaction over a commercially avail-
able catalyst in a fixed-bed reactor.

Description: The ConocoPhillips Hydrisom Process is a once-through re- ��

��
action and, for typical cat cracker streams, requires no recycle or cooling. ��
Hydrogen is added downstream of the olefin feed pump on ratio control ��
and the feed mixture is preheated by exchange with the fractionator
��
bottoms and/or low-pressure steam. The feed then flows downward
��
over a fixed bed of commercial catalyst. ��

The reaction is liquid-phase, at a pressure just above the bubble ��

point of the hydrocarbon/hydrogen mixture. The rise in reactor tem- ��
perature is a function of the quantity of butadiene in the feed and the
amount of butene saturation that occurs.
The Hydrisom Process can also be configured using a proprietary
catalyst to upgrade streams containing diolefins up to 50% or more,
e.g., steam cracker C4 steams, producing olefin-rich streams for use as
chemical, etherification and/or alkylation feedstocks. Installation: Ten units licensed worldwide, including an installation at
Installation of a Hydrisom unit upstream of an etherification and/ the ConocoPhillips Sweeny, Texas, Refinery.
or alkylation unit can result in a very quick payout of the investment
due to: Licensor: ConocoPhillips.
• Improved etherification unit operations
• Increased ether production
• Increased alkylate octane number
• Increased alkylate yield
• Reduced chemical and HF acid costs
• Reduced ASO handling
• Reduced alkylation unit utilities
• Upgraded steam cracker or other high diolefin streams (30% to
50%) for further processing.
Isomerization
Application: The widely used Butamer process is a high-efficiency, cost ��
��
effective means of meeting the demands for the production of isobu-
tane by isomerizing normal butane (nC4) to isobutane (i C4).
Motor-fuel alkylate is one blending component that has seen a sub-
stantial increase in demand because of its paraffinic, high-octane, low- ��
vapor pressure blending properties. Isobutane is a primary feedstock for
��
producing motor-fuel alkylate.
��
Description: UOP’s innovative hydrogen-once-through (HOT) Butamer pro-
cess results in substantial savings in capital equipment and utility costs by
eliminating the need for a product separator or recycle-gas compressor. ��
Typically, two reactors, in series flow, are used to achieve high
��
onstream efficiency. The catalyst can be replaced in one reactor while
operation continues in the other. The stabilizer separates the light gas ��
��
from the reactor effluent.
A Butamer unit can be integrated with an alkylation unit. In this
application, the Butamer unit feed is a side-cut from an isostripper col-
umn, and the stabilized isomerate is returned to the isostripper column.
Unconverted n-butane is recycled to the Butamer unit, along with n-bu-
tane from the fresh feed. Virtually complete conversion of n-butane to
isobutane can be achieved.
Feed: The best feeds for a Butamer unit contain the highest practical
n-butane content, and only small amounts of isobutane, pentanes and
heavier material. Natural gas liquids (NGL) from a UOP NGL recovery unit
can be processed in a Butamer unit.
Yield: The stabilized isomerate is a near-equilibrium mixture of isobutane
and n-butane with small amounts of heavier material. The light-ends
yield from cracking is less than 1 wt% of the butane feed.
Installation: More than 70 Butamer units have been commissioned, and
additional units are in design or construction. Butamer unit feed capaci-
ties range from 800 to 35,000+ bpsd (74 to 3,250 tpd).
Licensor: UOP LLC.
Isomerization ��

Application: The Par-Isom process is an innovative application using

high-performance nonchlorided-alumina catalysts for light-naphtha ��
isomerization. The process uses PI-242 catalyst, which approaches the
activity of chlorided alumina catalysts without requiring organic chloride
injection. The catalyst is regenerable and is sulfur and water tolerant.

Description: The fresh C5/ C6 feed is combined with make-up and re- ��

cycle hydrogen which is directed to a charge heater, where the reactants ��

are heated to reaction temperature. The heated combined feed is then ��
��
sent to the reactor. Either one or two reactors can be used in series, de-
pending on the speciﬁc application.
The reactor efﬂuent is cooled and sent to a product separator where
the recycle hydrogen is separated from the other products. Recovered ��
recycle hydrogen is directed to the recycle compressor and back to the
reaction section. Liquid product is sent to a stabilizer column where light ��
ends and any dissolved hydrogen are removed. The stabilized isomerate
product can be sent directly to gasoline blending.

Feed: Typical feed sources for the Par-Isom process include hydrotreated
light straight-run naphtha, light natural gasoline or condensate and light Installation: The first commercial Par-Isom process unit was placed in
raffinate from benzene extraction units. operation in 1996. There are currently 10 units in operation. The first
Water and oxygenates at concentrations of typical hydrotreated commercial application of PI-242 catalyst was in 2003, and the unit has
naphtha are not detrimental, although free water in the feedstock must demonstrated successful performance meeting all expectations.
be avoided. Sulfur suppresses activity, as expected, for any noble-metal
based catalyst. However, the suppression effect is fully reversible by sub- Licensor: UOP LLC.
sequent processing with clean feedstocks.

Yield: Typical product C5+ yields are 97 wt% of the fresh feed. The
product octane is 81 to 87, depending on the ﬂow conﬁguration and
feedstock qualities.
Isomerization
Application: Most of the implemented legislation requires limiting ��
benzene concentration in the gasoline pool. This has increased the ��
demand for high-performance C5 and C6 naphtha isomerization tech-
��
nology because of its ability to reduce the benzene concentration in ��
the gasoline pool while maintaining or increasing the pool octane. The
Penex process has served as the primary isomerization technology for ��
upgrading C5 / C6 light straight-run naphtha.

Description: UOP’s innovative hydrogen-once-through (HOT) Penex pro-

cess results in sustantial savings in capital equipment and utility costs by
eliminating the need for a product separator or recycle-gas compressor.
The Penex process is a fixed-bed process that uses high-activity chlo- ��
ride-promoted catalysts to isomerize C5 / C6 paraffins to higher-octane-
��
branched components. The reaction conditions promote isomerization ��
and minimize hydrocracking.
Typically, two reactors, in series flow, are used to achieve high
onstream efficiency. The catalyst can be replaced in one reactor while
operation continues in the other. The stabilizer separates light gas from
the reactor effluent.
Feed: Penex process can process feeds with high levels of C6 cyclics and
Products: For typical C5 / C6 feeds, equilibrium will limit the product to C7 components. In addition, feeds with substantial levels of benzene
83 to 86 RONC on a single hydrocarbon pass basis. To achieve higher can be processed without the need for a separate saturation section.
octane, UOP offers several schemes in which lower octane components
are separated from the reactor effluent and recycled back to the reac- Installation: UOP is the leading world-wide provider of isomerization tech-
tors. These recycle modes of operation can lead to product octane as nology. More than 120 Penex units are in operation. Capacities range
high as 93 RONC, depending on feed quality. from 1,000 bpsd to more than 25,000 bpsd of fresh feed capacity.
Yields: Licensor: UOP LLC.
Penex process: Octane 86
Penex process/DIH: Octane 90
Penex process/Molex process: Octane 91
DIP/Penex process/DIH: Octane 93
Isooctene/isooctane ��
��
��
��
��
��
� �
Application: New processes, RHT-isooctene and RHT-isooctane, can be
� ��
used to revamp existing MTBE units to isooctene/isooctane production. ��
��
Feeds include C4 iso-olefin feed from FCC, steam crackers, thermal �
� �
crackers or on-purpose iso-butylene from dehydrogenation units. The ��
��
processes uses a unique configuration for dimerization. A new selectiva-
�
tor is used, together with a dual-bed catalyst. �
��
The configuration is capable of revamping conventional or reactive �
��
distillation MTBE units. The process provides higher conversion, better ��
��
selectivity, conventional catalyst, and a new selectivator supports with ��
�
longer catalyst life with a dual catalyst application. ��
��
��
��
��
The process is designed to apply a hydrogenation unit to convert � �
��

isooctene into isooctane, if desired, by utilizing a dual-catalyst system, � ��

in the first and finishing reactors. The process operates at lower pressure ��
��
and provides lower costs for the hydrogenation unit. ��
�
Description: The feed is water washed to remove any basic compounds
that can poison the catalyst system. Most applications will be directed �
toward isooctene production. However as olefin specifications are re- ��
��
��
quired, the isooctene can be hydrogenated to isooctane, which is an
excellent gasoline blending stock.
The RHT isooctene process has a unique configuration; it is flexible
and can provide low per pass conversion through dilution, using a new provides the dilution, and reactor effluent is fed to the column at mul-
selectivator. The dual catalyst system also provides multiple advantages. tiple locations. Recycling does not increase column size due to the
The isobutylene conversion is 97–99 %, with better selectivity and yield unique configuration of the process. The isooctene is taken from the
together with enhanced catalyst life. The product is over 91% C8 olefins, debutanizer column bottom and is sent to OSBL after cooling or as is
and 5 – 9% C12 olefins, with very small amount of C16 olefins. sent to hydrogenation unit. The C4s are taken as overhead stream and
The feed after water wash, is mixed with recycle stream, which sent to OSBL or alkylation unit. Isooctene/product, octane (R+M)/2 is
provides the dilution (also some unreacted isobutylene) and is mixed expected to be about 105.
with a small amount of hydrogen. The feed is sent to the dual-bed re- If isooctane is to be produced the debutanizer bottom, isooctene
actor for isooctene reaction in which most of isobutylene is converted product is sent to hydrogenation unit. The isooctene is pumped to the
to isooctene and codimer. The residual conversion is done with single- required pressure (which is much lower than conventional processes),
resin catalyst via a side reactor. The feed to the side reactor is taken as mixed with recycle stream and hydrogen and is heated to the reaction
a side draw from the column and does contain unreacted isobutylene,
selectivator, normal olefins and non-reactive C4s. The recycle stream Continued 
Isooctene/Isooctane, continued Economics:
Isooctene Isooctane1
temperature before sending it the first hydrogenation reactor. This reac- CAPEX ISBL, MM USD
tor uses a nickel (Ni) or palladium (Pd) catalyst. (US Gulf Coast 1Q 06, 1,000 bpd) 8.15 5.5
If feed is coming directly from the isooctene unit, only a start-up Utilities basis 1,000-bpd isooctene/isooctane
heater is required. The reactor effluent is flashed, and the vent is sent Power, kWh 65 105
to OSBL. The liquid stream is recycled to the reactor after cooling (to Water, cooling, m /h 3 154 243
remove heat of reaction) and a portion is forwarded to the finishing Steam, HP, kg/h 3,870 4,650
reactor—which also applies a Ni or Pd catalyst (preferably Pd catalyst) — Basis: FCC feed (about 15 –20% isobutelene in C4 mixed stream)
and residual hydrogenation to isooctane reaction occurs. The isooctane 1These utilities are for isooctene / isooctane cumulative.
product, octane (R+M)/2 is >98.
The reaction occurs in liquid phase or two phase (preferably two Installation: Technology is ready for commercial application.
phases), which results in lower pressure option. The olefins in isooctene
product are hydrogenated to over 99%. The finishing reactor effluent is Licensor: Refining Hydrocarbon Technologies LLC.
sent to isooctane stripper, which removes all light ends, and the product
is stabilized and can be stored easily.
Lube and wax processing
Application: Vacuum gas oils (VGOs) are simultaneously extracted and ��
��
dewaxed on a single unit to produce low-pour aromatic extracts and
lube-base stocks having low-pour points. Low viscosity grades (60 SUS)
to bright stocks can be produced. With additional stages of filtration,
waxes can be deoiled to produce fully refined paraffin waxes.

Products: Lube-base stocks having low pour points (– 20°C). Very low
��
pour point aromatic extracts. Slack waxes or low-oil content waxes.
��
Description: Process Dynamics’ integrated extraction/dewaxing technol-
ogy is a revolutionary process combining solvent extraction and solvent
dewaxing onto a single unit, using a common solvent system, for extrac-
tion and dewaxing steps. This process offers the advantage of operating
a single unit rather than separate extraction and dewaxing units; thus, ��
reducing both capitol and operating costs. The more selective solvent
system produces lube-base stocks of higher quality and higher yields
when compared to other technologies.
Primary solvent and warm feed are mixed together; temperature is
controlled by adding the cosolvent solvent. Filtrate or wax may be re-
cycled for solids adjustment. Cold cosolvent is added, and the slurry is Installation: Basic engineering package for the first commercial unit has
filtered (or separated by other means). Solvents are recovered from the been completed.
primary filtrate producing an aromatic extract. The wax cake is repulped
with additional solvent/cosolvent mix and refiltered. Solvents are recov- Licensor: P. D. Licensing, LLC (Process Dynamics, Inc.).
ered from the filtrate producing a lube-base stock.
Extraction/dewaxing comparisons of 90 SUS stock
Furfural/MEK Process Dynamics
A B
Raffinate yields, vol% 53 60 71
Dewaxed oil properties:
Viscosity @40°C, cSt 16.5 18.7 20
Viscosity index 92 98 92
Pour pt., °F 5 5 5
Lube extraction
Application: Bechtel’s MP Refining process is a solvent-extraction pro-
�
cess that uses N-methyl-2-pyrrolidone (NMP) as the solvent to selec- ��
tively remove the undesirable components of low-quality lubrication oil,
��
which are naturally present in crude oil distillate and residual stocks. �
The unit produces paraffinic or naphthenic raffinates suitable for further
processing into lube-base stocks. This process selectively removes aro- ��
matics and compounds containing heteroatoms (e.g., oxygen, nitrogen ��
and sulfur).
��
Products: A raffinate that may be dewaxed to produce a high-qual-
��
ity lube-base oil, characterized by high viscosity index, good thermal
and oxidation stability, light color and excellent additive response. The
byproduct extracts, being high in aromatic content, can be used, in
some cases, for carbon black feedstocks, rubber extender oils and other
nonlube applications where this feature is desirable. Economics:
Description: The distillate or residual feedstock and solvent are contact- Investment (Basis: 10,000-bpsd feedrate
ed in the extraction tower (1) at controlled temperatures and flowrates capacity, 2006 US Gulf Coast), $/bpsd 3,000
required for optimum countercurrent, liquid-liquid extraction of the Utilities, typical per bbl feed:
feedstock. The extract stream, containing the bulk of the solvent, exits Fuel, 103 Btu (absorbed) 100
the bottom of the extraction tower. It is routed to a recovery section Electricity, kWh 2
to remove solvent contained in this stream. Solvent is separated from Steam, lb 5
the extract oil by multiple-effect evaporation (2) at various pressures, Water, cooling (25°F rise), gal 600
followed by vacuum flashing and steam stripping (3) under vacuum. Installation: This process is being used in 15 licensed units to produce high-
The raffinate stream exits the overhead of the extraction tower and is quality lubricating oils. Of this number, eight are units converted from phe-
routed to a recovery section to remove the NMP solvent contained in nol or furfural, with another three units under license for conversion.
this stream by flashing and steam stripping (4) under vacuum.
Overhead vapors from the steam strippers are condensed and com- Licensor: Bechtel Corp.
bined with solvent condensate from the recovery sections and are dis-
tilled at low pressure to remove water from the solvent (5). Solvent is
recovered in a single tower because NMP does not form an azeotrope
with water, as does furfural. The water is drained to the oily-water sew-
er. The solvent is cooled and recycled to the extraction section.
Continued 
Lube extraction
Application: Bechtel’s Furfural Refining process is a solvent-extraction
process that uses furfural as the solvent to selectively remove undesir-
�
able components of low lubrication oil quality, which are naturally pres- ��
ent in crude oil distillate and residual stocks. This process selectively re-
��
moves aromatics and compounds containing heteroatoms (e.g., oxygen, �
nitrogen and sulfur). The unit produces paraffinic raffinates suitable for ��
further processing into lube base stocks. ��
��
Products: A raffinate that may be dewaxed to produce a high-qual- ��
ity lube-base oil, characterized by high viscosity index, good thermal ��
and oxidation stability, light color and excellent additive response. The
��
byproduct extracts, being high in aromatic content, can be used, in
some cases, for carbon black feedstocks, rubber extender oils and other
nonlube applications where this feature is desirable.

Description: The distillate or residual feedstock and solvent are contact-

ed in the extraction tower (1) at controlled temperatures and flowrates
required for optimum countercurrent, liquid-liquid extraction of the
feedstock. The extract stream, containing the bulk of the solvent, exits Economics:
the bottom of the extraction tower. It is routed to a recovery section to Investment (Basis: 10,000-bpsd feed rate capacity,
remove solvent contained in this stream. Solvent is separated from the 2006 US Gulf Coast), $/ bpsd 3,100
extract oil by multiple-effect evaporation (2) at various pressures, fol- Utilities, typical per bbl feed:
lowed by vacuum flashing and steam stripping (3) under vacuum. The Fuel, 103 Btu (absorbed) 120
raffinate stream exits the overhead of the extraction tower and is routed Electricity, kWh 2
to a recovery section to remove the furfural solvent contained in this Steam, lb 5
stream by flashing and steam stripping (4) under vacuum. Water, cooling (25°F rise), gal 650
The solvent is cooled and recycled to the extraction section. Over-
Installation: For almost 60 years, this process has been or is being used
head vapors from the steam strippers are condensed and combined with
in over 100 licensed units to produce high-quality lubricating oils.
the solvent condensate from the recovery sections and are distilled at
low pressure to remove water from the solvent. Furfural forms an azeo- Licensor: Bechtel Corp.
trope with water and requires two fractionators. One fractionator (5)
separates the furfural from the azeotrope, and the second (6) separates
water from the azeotrope. The water drains to the oily-water sewer. The
solvent is cooled and recycled to the extraction section.
Lube hydrotreating
Application: The Bechtel Hy-Finishing process is a specialized hydrotreating
technology to remove impurities and improve the quality of paraffinic ��
and naphthenic lubricating base oils. In the normal configuration, the
hydrogen finishing unit is located in the processing scheme between the �
solvent extraction and solvent dewaxing units for a lube plant operating �
on an approved lube crude. In this application, the unit operates under
mild hydrotreating conditions to improve color and stability, to reduce ��
sulfur, nitrogen, oxygen and aromatics, and to remove metals. ��
��

Another application is Hy-Starting, which is a more severe hydro- �

treating process (higher pressure and lower space velocity) and upgrades
��
distillates from lower-quality crudes. This unit is usually placed before
solvent extraction in the processing sequence to upgrade distillate qual-
ity and, thus, improve extraction yields at the same rafﬁnate quality. ��

Description: Hydrocarbon feed is mixed with hydrogen (recycle plus

makeup), preheated, and charged to a fixed-bed hydrotreating reac-
tor (1). Reactor effluent is cooled in exchange with the mixed feed-hy-
drogen stream. Gas-liquid separation of the effluent occurs first in the
hot separator (2) then in the cold separator (3). The hydrocarbon liquid
stream from each of the two separators is sent to the product stripper (4)
to remove the remaining gas and unstabilized distillate from the lube-oil
product. The product is then dried in a vacuum flash (5). Gas from the
cold separator is amine-scrubbed (6) to remove H2S before compression
in the recycle hydrogen compressor (7).
Economics:
Investment (Basis 7,000-bpsd feedrate capacity,
2006 US Gulf Coast), $/bpsd 5,100
Utilities, typical per bbl feed:
Fuel, 103 Btu (absorbed) 20
Electricity, kWh 5
Steam, lb 15
Water, cooling (25°F rise), gal 400
Licensor: Bechtel Corp.
Lube hydrotreating
Application: Hy-Raff is a new process to hydrotreat raffinates from an ��
extraction unit of a solvent-based lube oil plant for upgrading standard
Group I lube-base oils to produce Group II base oils. Sulfur is reduced to
below 0.03 wt% and saturates are increased to greater than 90 wt%. �
The integration of this process into an existing base oil plant allows �
the operator to cost-effectively upgrade base-oil products to the new
specifications rather than scrapping the existing plant and building an � ��
��
expensive new hydrocracker-based plant. �
��
The product from the Hy-Raff unit is a lube-base oil of sufficient
quality to meet Group II specifications. The color of the finished prod- � ��
uct is significantly improved over standard-base oils. Middle distillate
byproducts are of sufficient quality for blending into diesel. � �
��
Description: Raffinate feed is mixed with hydrogen (recycle plus make- ��
up), preheated, and charged to a fixed-bed hydrotreating reactor (1).
The reactor effluent is cooled in exchange with the mixed feed-hydrogen
stream. Gas-liquid separation of the effluent occurs first in the hot sepa-
rator (2) then in the cold separator (3). The hydrocarbon liquid stream
from each of the two separators is sent to the product stripper (4) to
remove the remaining gas and unstabilized distillate from the lube-oil
product, and product is dried in a vacuum flash (5). Gas from the cold
separator is amine-scrubbed (6) for removal of H2S before compression
in the recycle-hydrogen compressor (7).

Economics:
Investment (Basis 7,000-bpsd feedrate capacity,
2006 U.S. Gulf Coast), $/bpsd 7,100
Utilitiies, typical per bbl feed:
Fuel, 103 Btu (absorbed) 70
Electricity, kWh 5
Steam, lb 15
Water, cooling (25°F rise), gal 200

Licensor: Bechtel Corp.

Lube oil refining, spent
Application: The Revivoil process can be used to produce high yields of
��
premium quality lube bases from spent motor oils. Requiring neither
acid nor clay treatment steps, the process can eliminate environmental ��
and logistical problems of waste handling and disposal associated with ��
��
conventional re-refining schemes. ��
��
Description: Spent oil is distilled in an atmospheric flash distillation ��
column to remove water and gasoline and then sent to the Thermal ��
��
Deasphalting (TDA) vacuum column for recovery of gas oil overhead
and oil bases as side streams. The energy-efficient TDA column features
��
excellent performance with no plugging and no moving parts. Metals ��
and metalloids concentrate in the residue, which is sent to an optional
��
Selectopropane unit for brightstock and asphalt recovery. This scheme ��
��
is different from those for which the entire vacuum column feed goes
through a deasphalting step; Revivoil’s energy savings are significant,
and the overall lube oil base recovery is maximized. The results are sub-
stantial improvements in selectivity, quality and yields.
The final, but very important step for base oil quality is a specific
hydrofinishing process that reduces or removes remaining metals and pounds in the final product at a level inferior to 5 wppm (less than 1
metalloids, Conradson Carbon, organic acids, and compounds contain- wt% PCA - IP346 method).
ing chlorine, sulfur and nitrogen. Color, UV and thermal stability are
Economics: The process can be installed stepwise or entirely. A simpler
restored and polynuclear aromatics are reduced to values far below the
scheme consists of the atmospheric flash, TDA and hydrofinishing unit
latest health thresholds. Viscosity index remains equal to or better than
and enables 70 – 80% recovery of lube oil bases. The Selectopropane
the original feed. For metal removal (> 96%) and refining-purification
unit can be added at a later stage, to bring the oil recovery to the 95%
duty, the multicomponent catalyst system is the industry’s best.
level on dry basis. Economics below show that for two plants of equal
Product quality: The oil bases are premium products; all lube oil base capacity, payout times before taxes are two years in both cases.
specifications are met by Revivoil processing from Group 1 through
Investment: Basis 100,000 metric tpy, water-free, ISBL 2004 Gulf
Group 2 of the API basestocks definitions. Besides, a diesel can be ob-
Coast, million US$
tained, in compliance with the EURO 5 requirements (low sulfur).
Configuration 1 (Atm. flash, TDA and Hydrofinishing units) 30
Health & safety and environment: The high-pressure process is in line Configuration 2 (Same as above + Selectopropane unit) 35
with future European specifications concerning carcinogenic PNA com-

Continued 
Lube oil refining, spent, continued
Utilities: Basis one metric ton of water-free feedstock
Config. 1 Config. 2
Electrical power, kWh 45 55
Fuel, million kcal 0.62 0.72
Steam, LP, kg — 23.2
Steam, MP, kg 872 890
Water, cooling, m3 54 59

Installation: Ten units have been licensed using all or part of the Revivoil
Technology.

Licensor: Axens and Viscolube SpA.

Lube treating
��
Application: Lube raffinates from extraction are dewaxed to provide base- ��
stocks having low pour points (as low as –35°C). Basestocks range from ��
��
��
light stocks (60N) to higher viscosity grades (600N and bright stock). ��
��
Byproduct waxes can also be upgraded for use in food applications. ��
��
��
Feeds: DILCHILL dewaxing can be used for a wide range of stocks that boil ��
above 550°F, from 60N up through bright stock. In addition to raffinates ��
from extraction, DILCHILL dewaxing can be applied to hydrocracked stocks
and to other stocks from raffinate hydroconversion processes. ��
��
Processes: Lube basestocks have low pour points. Although slack wax-
��
es containing 2–10 wt% residual oil are the typical byproducts, lower- ��
oil-content waxes can be produced by using additional dewaxing and/or
“warm-up deoiling” stages. ��

Description: DILCHILL is a novel dewaxing technology in which wax crystals

are formed by cooling waxy oil stocks, which have been diluted with ketone
solvents, in a proprietary crystallizer tower that has a number of mixing
stages. This nucleation environment provides crystals that filter more quickly Economics: Depend on the slate of stocks to be dewaxed, the pour
and retain less oil. This technology has the following advantages over con- point targets and the required oil-in-wax content.
ventional incremental dilution dewaxing in scraped-surface exchangers: less
filter area is required, less washing of the filter cake to achieve the same Utilities: Depend on the slate of stocks to be dewaxed, the pour point
oil-in-wax content is required, refrigeration duty is lower, and only scraped targets and the required oil-in-wax content.
surface chillers are required which means that unit maintenance costs are Installation: The first application of DILCHILL dewaxing was the conver-
lower. No wax recrystallization is required for deoiling. sion of an ExxonMobil affiliate unit on the U.S. Gulf Coast in 1972. Since
Warm waxy feed is cooled in a prechiller before it enters the DILCHILL that time, 10 other applications have been constructed. These applica-
crystallizer tower. Chilled solvent is then added in the crystallizer tower tions include both grassroots units and conversions of incremental dilu-
under highly agitated conditions. Most of the crystallization occurs in tion plants. Six applications use “warm-up deoiling.”
the crystallizer tower. The slurry of wax/oil/ketone is further cooled in
scraped-surface chillers and the slurry is then filtered in rotary vacuum Licensor: ExxonMobil Research and Engineering Co.
filters. Flashing and stripping of products recover solvent. Additional fil-
tration stages can be added to recover additional oil or/to produce low-
oil content saleable waxes.
Lube treating
��
Application: Process to produce lube oil raffinates with high viscosity ��
index from vacuum distillates and deasphalted oil. ��
��
��
Feeds: Vacuum distillate lube cuts and deasphalted oils. ��

Products: Lube oil rafﬁnates of high viscosity indices. The rafﬁnates con- ��
��
��
tain substantially all of the desirable lubricating oil components present ��
in the feedstock. The extract contains a concentrate of aromatics that ��
��
may be utilized as rubber oil or cracker feed. ��
��

Description: This liquid-liquid extraction process uses furfural or N-

��
methyl pyrrolidone (NMP) as the selective solvent to remove aromat- ��
ics and other impurities present in the distillates and deasphalted oils. ��
��
The solvents have a high solvent power for those components that are
��
unstable to oxygen as well as for other undesirable materials includ- ��
��
ing color bodies, resins, carbon-forming constituents and sulfur com- ��

pounds. In the extraction tower, the feed oil is introduced below the top
at a predetermined temperature. The raffinate phase leaves at the top
of the tower, and the extract, which contains the bulk of the furfural, is
withdrawn from the bottom. The extract phase is cooled and a so-called
“pseudo raffinate“ may be sent back to the extraction tower. Multi-
stage solvent recovery systems for raffinate and extract solutions secure
energy efficient operation.
Utility requirements (typical, Middle East Crude), units per m3 of feed:
Electricity, kWh 10
Steam, MP, kg 10
Steam, LP, kg 35
Fuel oil, kg 20
Water, cooling, m3 20
Installation: Numerous installations using the Uhde (Edeleanu) propri-
etary technology are in operation worldwide. The most recent is a com-
plete lube-oil production facility licensed to the state of Turkmenistan.
Licensor: Uhde GmbH.
Mercaptan removal
��
Application: Extraction of mercaptans from gases, LPG, lower boiling
fractions and gasolines, or sweetening of gasoline, jet fuel and diesel by ��
in situ conversion of mercaptans into disulfides.
�
Products: Essentially mercaptan sulfur-free, i.e., less than 5 ppmw, and
concomitant reduced total sulfur content when treated by Merox ex- �
��
��
traction technique.

Description: Merox units are designed in several ﬂow conﬁgurations, ��

depending on feedstock type and processing objectives. All are charac- ��
��
terized by low capital and operating costs, ease of operation and mini- ��
mal operator attention.
��
��
Extraction: Gases, LPG and light naphtha are countercurrently extracted
(1) with caustic containing Merox catalyst. Mercaptans in the rich caustic
are oxidized (2) with air to disulfides that are decanted (3) before the
regenerated caustic is recycled.
Heavy
Sweetening: Minalk is now the most prevalent Merox gasoline sweeten- naphtha
ing scheme. Conversion of mercaptans into disulfides is accomplished
Product Gas LPG Gasoline fixed bed
with a fixed bed of Merox catalyst that uses air and a continuous injec-
Scheme Ext. Ext. Minalk
tion of only minute amounts of alkali. Sweetened gasoline from the 3
Est. plant capital, modular $10 2,600 2,000 1,100 1,500
reactor typically contains less than one ppm sodium.
Direct operating cost, ¢/bbl
Heavy gasoline, condensate, kerosine/jet fuel and diesel can be
(¢/10 6 scf) (1.5) 0.4 0.2 1.0
sweetened in a fixed-bed unit that closely resembles Minalk, except that
a larger amount of more concentrated caustic is recirculated intermit- Installations: Capacity installed and under construction exceeds 13 million
tently over the catalyst bed. A new additive, Merox Plus catalyst activa- bpsd. More than 1,600 units have been commissioned to date, with capaci-
tor, can be used to greatly extend catalyst life. ties between 40 and 140,000 bpsd. UOP has licensed gas Merox extraction
units with capacities as high as 2.9-billion-scfd for mercaptan control.
Economics: Typical capital investment and operating costs of some
Merox process schemes are given based on 2004 dollars for a 10,000- Licensor: UOP LLC.
bpsd capacity liquid unit and 10 million-scfd gas unit with modular de-
sign and construction.
NOx abatement
Application: Flue gases are treated with ammonia via ExxonMobil’s propri-
��
etary selective noncatalytic NOx reduction technology—THERMAL DeNOx.
��
NOx plus ammonia (NH3) are converted to elemental nitrogen and water ��
��
if temperature and residence time are appropriate. The technology has ��
been widely applied since it was first commercialized in 1974.

Products: If conditions are appropriate, the ﬂue gas is treated to achieve ��
NOx reductions of 40% to 70%+ with minimal NH3 slip or leakage. ��
��
Description: The technology involves the gas-phase reaction of NO with ��
��
NH3 (either aqueous or anhydrous) to produce elemental nitrogen if con-
��
ditions are favorable. Ammonia is injected into the ﬂue gas using steam
or air as a carrier gas into a zone where the temperature is 1,600°F ��
to 2,000°F. This range can be extended down to 1,300°F with a small ��
amount of hydrogen added to the injected gas. For most applications,
wall injectors are used for simplicity of operation.

Yield: Cleaned ﬂue gas with 40% to 70%+ NOx reduction and less than
10-ppm NH3 slip.

Economics: Considerably less costly than catalytic systems but relatively McIntyre, A. D., “Applications of the THERMAL DeNOx process to
variable depending on scale and site speciﬁcs. Third-party studies have FBC boilers,” CIBO 13th Annual Fluidized Bed Conference, Lake Charles,
estimated the all-in cost at about 600 US$ / ton of NOx removed. Louisiana, 1997.

Installation: Over 135 applications on all types of ﬁred heaters, boilers Licensor: ExxonMobil Research and Engineering Co., via an alliance with
and incinerators with a wide variety of fuels (gas, oil, coal, coke, wood Engineers India Ltd. (for India) and Hamon Research-Cottrell (for the rest
and waste). The technology can also be applied to full-burn FCCU re- of the world).
generators.

Reference: McIntyre, A. D., “Applications of the THERMAL DeNOx pro-

cess to utility and independent power production boilers,” ASME Joint
International Power Generation Conference, Phoenix, 1994.
McIntyre, A. D., “The THERMAL DeNOx process: Liquid fuels appli-
cations,” International Flame Research Foundation’s 11th Topic Oriented
Technical Meeting, Biarritz, France, 1995.
NOx reduction, low-temperature
Application: The LoTOx low-temperature oxidation process removes NOx
from flue gases in conjunction with BELCO’s EDV wet scrubbing system. ��
Ozone is a very selective oxidizing agent; it converts relatively insoluble
NO and NO2 to higher, more soluble nitrogen oxides. These oxides are
easily captured in a wet scrubber that is controlling sulfur compounds ��
and/or particulates simultaneously.
��
Description: In the LoTOx process, ozone is added to oxidize insoluble
NO and NO2 to highly oxidized, highly soluble species of NOx that can
��
be effectively removed by a variety of wet or semi-dry scrubbers. Ozone, ��
a highly effective oxidizing agent, is produced onsite and on demand
by passing oxygen through an ozone generator—an electric corona de-
vice with no moving parts. The rapid reaction rate of ozone with NOx
results in high selectivity for NOx over other components within the gas
stream.
Thus, the NOx in the gas phase is converted to soluble ionic com- the past seven years, winning the prestigious 2001 Kirkpatrick Chemical
pounds in the aqueous phase; the reaction is driven to completion, thus Engineering Technology Award. Currently, there are eight commercial
removing NOx with no secondary gaseous pollutants. The ozone is con- applications, including boilers and FCC units. Many other EDV scrubbers
sumed by the process or destroyed within the system scrubber. All sys- have been designed for LoTOx application. Pilot-scale demonstrations
tem components are proven, well-understood technologies with a his- have been completed on coal- and petroleum-coke fired boilers, as well
tory of safe and reliable performance. as refinery FCC units.

Operating conditions: Ozone injection typically occurs in the ﬂue-gas Reference: Confuorto, et al., “LoTOx technology demonstration at Mar-
stream upstream of the scrubber, near atmospheric pressure and at athon Ashland Petroleum LLC’s reﬁnery at Texas City, Texas,” NPRA An-
temperatures up to roughly 150°C. For higher-temperature streams, the nual Meeting, March 2004, San Antonio.
ozone is injected after a quench section of the scrubber, at adiabatic
Licensor: Belco Technologies Corp., as a sub-licensor for The BOC
saturation, typically 60°C to 75°C. High-particulate saturated gas and
Group, Inc.
sulfur loading (SOx or TRS) do not cause problems.

Economics: The costs for NOx control using this technology are espe-
cially low when used as a part of a multi-pollutant control scenario.
Sulfurous and particulate-laden streams can be treated attractively as no
pretreatment is required by the LoTOx system.

Installation: The technology has been developed and commercialized over

Oxygen enrichment for Claus units Liquid oxygen tank
Claus plant process
Application: Increase the capacity of Claus plants and decompose haz- Vaporizer control system
ardous materials such as ammonia
Controller
Description: As “clean fuels” regulations become effective, refiners
must recover more sulfur in their Claus plants. As a byproduct of deep Measuring and
1 control unit Steam
desulfurization, ammonia is generated and typically must be decom- FLOWTRAIN
posed in the Claus plant. To upgrade the sulfur recovery units (SRUs) 2 4 Process gas
accordingly, oxygen enrichment is an effcient and low-cost option. to catalytic
Air reactors
Oxygen enrichment can increase sulfur capacity substantially and is Onsite ASU
capable of decomposing ammonia from sour-water stripper gas very 3
efficiently.
Oxygen addition can be done in three levels, depending on the Oxygen pipeline Acid gas plus sour BFW
water stripper gas
required capacity increase:
1 Alternative oxygen sources
1. Up to approximately 28% oxygen. Oxygen is simply added to 2 FLOWTRAIN with all required safety features
the Claus furnace air. This can raise sulfur capacity by up to 25%. 3 Oxygen injection and mixing device
4 Claus reaction furnace with burner for air and/or oxygen enriched operation
2. Up to approximately 40% oxygen. The burner of the Claus fur-
nace must be replaced. Up to 50% additional sulfur capacity can be
achieved by this method.
3. Beyond 40% oxygen. The temperature in the Claus furnace is
elevated so high that the product gas must be recycled to maintain for a short and highly turbulent flame, which ensures good approach
temperature control. This process is expensive and, therefore, rarely toward equilibrium for Claus operation and for the decomposition of
applied. ammonia.
Oxygen sources can be liquid oxygen tanks, onsite ASUs or pipe-
Economics: As oxygen enrichment provides substantial additional
line supply. Oxygen consumption in Claus plants fluctuates widely in
Claus capacity, it is a low-cost alternative to building an additional Claus
most cases; thus, tanks are the best choice due to ease of operation,
plant. It can save investment, manpower and maintenance. Installed
flexibility and economy. For oxygen addition into the CS air duct, a
cost for oxygen enrichment per level 1 is typically below $250,000.
number of safety rules must be observed. The oxygen metering device
For level 2, the investment costs range from $200,000 to $500,000
FLOWTRAIN contains all of the necessary safety features, including
and depend on the size of the Claus plant. Operating costs are varied
flow control, low-temperature and low-pressure alarm and switch-off,
and depend on the duration of oxygen usage. Typically, annual costs
and safe standby operation. All features are connected to the Claus
of oxygen enrichment are estimated as 10% to 40% of the cost for a
plants’ process control system.
Claus plant, providing the same additional sulfur capacity. Due to im-
An effcient mixing device ensures even oxygen distribution in the
proved ammonia destruction maintenance work, as cleaning of heat
Claus air. A proprietary Claus burner was developed especially for ap-
plication for air- and oxygen-enriched operations. This burner provides Continued 
Oxygen enrichment for Claus units, continued

exchanger tubes from ammonium salts and the respective corrosion

become substantially less.

Installations: Over 10, plus numerous test installations to quantify the

effects of capacity increase and ammonia decomposition.

Reference: Reinhardt, H. J. and P. M. Heisel, “Increasing the capacity of

Claus plants with oxygen,” Reports on Science and Technology, No. 61
(1999), p. 2.

Contributor: Linde AG, Division Gas and Engineering.

Oxygen enrichment for FCC units ��
��
��
Application: Increase the throughput capacity by up to 50% and/or con- ��
version in FCC units; process heavier feeds; overcome blower limita- ��
�
tions, also temporarily. � ��
�
��

Description: “Clean fuels” regulations are being implemented. Plus, the � �

��
� �
demand for transport fuels continually shifts toward more kerosine and
��
diesel. The reason is that the regulations and the change in demand are ��
totally independent developments. But both contribute to the require- �
�
ment of more flexibility in fluid catalytic cracking units (FCCUs). Conse-
quently, FCCUs require more flexibility to treat a wider range of feeds, ��
especially heavier feeds, and increasing throughput capacity. Both goals ��

can be achieved via oxygen enrichment in the FCC regeneration. ��
In the FCC reactor, long-chain hydrocarbons are cleaved into shorter ��
��
chains in a ﬂuidized-bed reactor at 450 –550°C. This reaction produces ��
coke as a byproduct that deposits on the catalyst. To remove the coke ��

from the catalyst, it is burned off at 650 –750°C in the regenerator. The
regenerated catalyst is returned to the reactor.
Oxygen enrichment, typically up to 27 vol% oxygen, intensifies cata- products, such as naphtha. Likewise, lower value products increased
lyst regeneration and can substantially raise throughput capacity and/or only 5%, as fuel gas. The net profit increased substantially. Installed cost
conversion of the FCC unit. Oxygen sources can be liquid oxygen tanks, for oxygen enrichment is typically below $250,000.
onsite ASUs or pipeline supply. Oxygen consumption in FCC units fluctu- Operating costs will depend on the cost for oxygen and the duration
ates widely in most cases; thus, tanks are the best choice with respect to of oxygen enrichment. Economical oxygen usage can be calculated on
ease of operation, flexibility and economy. a case-by-case basis and should include increased yields of higher-value
For oxygen addition into the CS air duct, a number of safety rules products and optional usage of lower-value feeds.
must be observed. The oxygen metering device FLOWTRAIN contains all
necessary safety features, including flow control, low-temperature and Installations: Currently, four units are in operation, plus test installations
low-pressure alarm and switch-off, and safe standby operation. All of to quantify the effects of higher capacity and conversion levels.
these features are connected to the FCC units’ process control system.
Reference: Heisel, M. P., C. Morén, A. Reichhold, A. Krause, A. Berlanga,
An efficient mixing device ensures even oxygen distribution in the air
“Cracking with Oxygen,” Linde Technology 1, 2004.
feed to the FCC regeneration.
Contributor: Linde AG, Division Gas and Engineering.
Economics: Oxygen enrichment in FCC regeneration is economically
favorable in many plants. For example, one refinery increased through-
put by 15%. The net improvement was a 26% increase in higher-value
Olefin etherification
Application: New processing methods improve etherification of C4– C7
reactive olefins including light catalytic naphtha (LCN) with alcohol (e.g., ��
��
��
��
��
��
��
��
��
��
methanol and ethanol). The processes, RHT-MixedEthers, RHT-MTBE, ��
RHT-ETBE, RHT-TAME and RHT-TAEE, use unique concepts to achieve the ��
��
��
maximum conversion without applying cumbersome catalyst in the col- ��
��
��
umn. The processing economics provide improvements over other avail-
�
able ether technologies currently available. The technology suite can ��
�
��
��
� ��
be applied to ethyl tertiary butyl ether (ETBE) production in which wet ��
��
ethanol can be used in place of dry ethanol. The drier can be eliminated,
�
which is approximately half the cost for an etherification unit. The RHT ��
��
��
ethers processes can provide the highest conversion with unique mul-
��
tiple equilibrium stages. ��

��
Description: The feed is water washed to remove basic compounds
that are poisons for the resin catalyst of the etherification reaction. The
C4 ethers—methyl tertiary butyl ether (MTBE)/ETBE), C5– tertiary amyl
methyl ether (TAME/ tertiary amyl ethyl ether (TAEE) and C6 /C7 ethers
are made in this process separately. The reaction is difficult; heavier
ethers conversion of the reactive olefins are equilibrium conversion of separate the ether and heavy hydrocarbons from C4 or C5 hydrocarbons,
about 97% for MTBE and 70% for TAME and much lower for C6/C7 which are taken as overhead. Single or multiple draw offs are taken from
ethers are expected. the fractionation column. In the fractionation column, unreacted olefins
Higher alcohols have similar effects (azeotrope hydrocarbon/alcohol (C4 or C5) are sent to the finishing reactor (5). This stream normally does
relationship decreases when using methanol over ethanol). The equilib- not require alcohol, since azeotrope levels are available. But, some ad-
rium conversions and azeotrope effects for higher ethers are lower, as is ditional alcohol is added for the equilibrium-stage reaction. Depending
expected. After the hydrocarbon feed is washed, it is mixed with alcohol on the liquid withdrawn (number of side draws), the conversion can be
with reactive olefin ratio control with alcohol. enhanced to a higher level than via other conventional or unconven-
The feed mixture is heated to reaction temperature (and mixed with tional processes.
recycle stream (for MTBE/ETBE only) and is sent to the first reactor (1), By installing multiple reactors, it is possible to extinct the olefins
where equilibrium conversion is done in the presence of sulfonated resin within the raffinate. The cost of side draws and reactors can achieve
catalyst, e.g. Amberlyst 15 or 35 or equivalent from other vendors. pay-off in 6 to 18 months by the higher catalyst cost as compared to
Major vaporization is detrimental to this reaction. Vapor-phase re- other processes. This process could provide 97– 99.9% isobutene con-
active olefins are not available for reaction. Additionally at higher tem- version in C4 feed (depending on the configuration) and 95 – 98+% of
peratures, there is slight thermal degradation of the catalyst occurs. The isoamylenes in C5 stream.
reactor effluent is sent to fractionator (debutanizer or depentanizer) to Continued 
catalyst. Distillation is done at optimum conditions. Much lower steam
Olefin etherification, continued consumption for alcohol recovery. For example, the C5 feed case requires
less alcohol with RHT configuration (azeotropic alcohol is not required)
The ether product is taken from the bottom, cooled and sent to the
and lowers lower steam consumption.
storage. The raffinate is washed in extractor column (6) with and is sent
to the OSBL. The water/alcohol mixture is sent to alcohol recovery col- Economics:
umn (7) where the alcohol is recovered and recycled as feed. CAPEX ISBL, MM USD (US Gulf Coast 1Q06,
For ETBE and TAEE, ethanol dehydration is required for most of the 1,000-bpd ether product) 9.1
processes, whereas for RHT process, wet ethanol can be used providing Utilities Basis 1,000 bpd ether
maximum conversions. If need be, the TBA specification can be met by Power kWh 45.0
optimum design with additional equipment providing high ETBE yield Water, cooling m3/ h 250
and conversion. Cost of ethanol dehydration is much more than the Steam MP, Kg / h 6,000
present configuration for the RHT wet-ethanol process. Basis: FCC Feed (about 15–20% isobutylene in C4 mixed stream)
The total capital cost /economics is lower with conventional catalyst
usage, compared to other technologies, which use complicated struc- Commercial units: Technology is ready for commercialization.
ture, require installing a manway (cumbersome) and require frequently
catalyst changes outs.
Licensor: Refining Hydrocarbon Technologies LLC.
The RHT ether processes can provide maximum conversion as com-
pared to other technologies with better economics. No complicated or
proprietary internals for the column including single source expensive
Olefins recovery
Application: Recover high-purity hydrogen (H2) and C2+ liquid products
from refinery offgases using cryogenics.
�
Description: Cryogenic separation of refinery offgases and purges con-
taining 10– 80% H2 and 15 – 40% hydrocarbon liquids such as ethylene, ��
��
ethane, propylene, propane and butanes. Refinery offgases are option-
ally compressed and then pretreated (1) to remove sulfur, carbon di-
oxide ( CO2), H2 O and other trace impurities. Treated feed is partially � �
condensed in an integrated multi-passage exchanger system (2) against �

returning products and refrigerant.

Separated liquids are sent to a demethanizer (3) for stabilization � ��
while hydrogen is concentrated (4) to 90 – 95%+ purity by further cool- ��
ing. Methane, other impurities, and unrecovered products are sent to � ��
��
��
fuel or optionally split into a synthetic natural gas (SNG) product and ��
low-Btu fuel. Refrigeration is provided by a closed-loop system (5). Mixed
C2+ liquids from the demethanizer can be further fractionated (6) into
finished petrochemical feeds and products such as ethane, ethylene,
propane and propylene.
and refinery fuel gas. Several process and refrigeration schemes used
Operating conditions: Feed capacities from 10 to 150+ million scfd. Feed since 1987 with the most recent plant startup in 2001.
pressures as low as 150 psig. Ethylene recoveries are greater than 95%,
with higher recoveries of ethane and heavier components. Hydrogen Reference: US Patents 6,266,977 and 6,560,989.
recoveries are better than 95% recovery. Trautmann, S. R. and R. A. Davis, “Refinery offgases—alternative
sources for ethylene recovery and integration,” AIChE Spring Meeting,
Economics: Hydrogen is economically co-produced with liquid hydro- New Orleans, March 14, 2002, Paper 102d.
carbon products, especially ethylene and propylene, whose high value
can subsidize the capital investment. High hydrocarbon liquid products Licensor: Air Products and Chemicals Inc.
recovery is achieved without the cost for feed compression and subse-
quent feed expansion to fuel pressure. Power consumption is a function
of hydrocarbon quantities in the feed and feed pressure. High-purity
hydrogen is produced without the investment for a “back-end” PSA
system. Project costs can have less than a two-year simple payback.

Installations: Five operating reﬁnery offgas cryogenic systems processing

FCC offgas, cat reformer offgas, hydrotreater purge gas, coker offgas
Olefins—butenes extractive distillation
Application: Separation of pure C4 olefins from olefinic/paraffinic C4 mix- ��
tures via extractive distillation using a selective solvent. BUTENEX is the
Uhde technology to separate light olefins from various C4 feedstocks,
which include ethylene cracker and FCC sources.

Description: In the extractive distillation (ED) process, a single-com- ��

pound solvent, N-Formylmorpholine (NFM), or NFM in a mixture with ��
��
further morpholine derivatives, alters the vapor pressure of the com- ��
ponents being separated. The vapor pressure of the olefins is lowered
more than that of the less soluble paraffins. Paraffinic vapors leave the
top of the ED column, and solvent with olefins leaves the bottom of the
ED column.
The bottom product of the ED column is fed to the stripper to ��
��
separate pure olefins (mixtures) from the solvent. After intensive heat
exchange, the lean solvent is recycled to the ED column. The solvent,
which can be either NFM or a mixture including NFM, perfectly satisfies
the solvent properties needed for this process, including high selectivity,
thermal stability and a suitable boiling point.

Economics:
Consumption per metric ton of FCC C4 fraction feedstock:
Steam, t / t 0.5 – 0.8
Water, cooling ( T = 10°C ), m3/ t 15.0
Electric power, kWh/t 25.0
Product purity:
n - Butene content 99.+ wt.– % min.
Solvent content 1 wt.– ppm max.

Installation: Two commercial plants for the recovery of n - butenes have

been installed since 1998.

Licensor: Uhde GmbH.

Olefins—dehydrogenation of light
paraffins to olefins ��
��
Application: The Uhde STeam Active Reforming (STAR) process produces ��
��
(a) propylene as feedstock for polypropylene, propylene oxide, cumene, ��
��
acrylonitrile or other propylene derivatives, and (b) butylenes as feed- ��
stock for methyl tertiary butyl ether (MTBE), alkylate, isooctane, polybu-
tylenes or other butylene derivatives. ��
��
Feed: Liquefied petroleum gas (LPG) from gas fields, gas condensate ��
fields and refineries.
��
Product: Propylene (polymer- or chemical-grade); isobutylene; n-butylenes; ��
��
��
high-purity hydrogen (H2) may also be produced as a byproduct. ��
��
Description: The fresh paraffin feedstock is combined with paraffin re- ��

cycle and internally generated steam. After preheating, the feed is sent
to the reaction section. This section consists of an externally fired tubular
fixed-bed reactor (Uhde reformer) connected in series with an adiabat-
ic fixed-bed oxyreactor (secondary reformer type). In the reformer, the
Apart from light-ends, which are internally used as fuel gas, the
endothermic dehydrogenation reaction takes place over a proprietary,
olefin is the only product. High-purity H2 may optionally be recovered
noble metal catalyst.
from light-ends in the gas separation section.
In the adiabatic oxyreactor, part of the hydrogen from the intermediate
product leaving the reformer is selectively converted with added oxygen Economics: Typical specific consumption figures (for polymer-grade
or air, thereby forming steam. This is followed by further dehydrogenation propylene production) are shown (per metric ton of propylene product,
over the same noble-metal catalyst. Exothermic selective H2 conversion including production of oxygen and all steam required):
in the oxyreactor increases olefin product space-time yield and supplies
Propane, kg/metric ton 1,200
heat for further endothermic dehydrogenation. The reaction takes place Fuel gas, GJ/metric ton 6.4
at temperatures between 500– 600°C and at 4 – 6 bar. Circul. cooling water, m3/metric ton 170
The Uhde reformer is top-fired and has a proprietary “cold” out- Electrical energy, kWh/metric ton 100
let manifold system to enhance reliability. Heat recovery utilizes process
heat for high-pressure steam generation, feed preheat and for heat re- Installation: Two commercial plants using the STAR process for dehydro-
quired in the fractionation section. genation of isobutane to isobutylene have been commissioned (in the
After cooling and condensate separation, the product is subse- US and Argentina). More than 60 Uhde reformers and 25 Uhde second-
quently compressed, light-ends are separated and the olefin product is ary reformers have been constructed worldwide.
separated from unconverted paraffins in the fractionation section. Continued 
Olefins, continued
References: Heinritz-Adrian, M., “Advanced technology for C3/C4 dehy-
drogenation, “ First Russian & CIS GasTechnology Conference, Moscow,
Russia, September 2004.
Heinritz-Adrian, M., N. Thiagarajan, S. Wenzel and H. Gehrke,
“STAR—Uhde’s dehydrogenation technology (an alternative route to
C3- and C4-olefins),” ERTC Petrochemical 2003, Paris, France, March
2003.
Thiagarajan, N., U. Ranke and F. Ennenbach, “Propane/butane de-
hydrogenation by steam active reforming,” Achema 2000, Frankfurt,
Germany, May 2000.

Licensor: Uhde GmbH.

Oligomerization—C3 /C4 cuts
Application: To dimerize light olefins such as ethylene, propylene and
butylenes using the Dimersol process. The main applications are: ��
��
• Dimerization of propylene, producing a high-octane, low-boiling
point gasoline called Dimate ��
• Dimerization of n-butylene producing C8 olefins for plasticizer syn-
��
thesis.
The C3 feeds are generally the propylene cuts from catalytic crack- ��
ing units. The C4 cut source is mainly the raffinate from butadiene and
� � � �
isobutylene extraction. ��

Description: Dimerization is achieved in the liquid phase at ambient

temperature by means of a soluble catalytic complex. One or several ��
��
��
reactors (1) in series are used. After elimination of catalyst (2, 3), the ��
products are separated in an appropriate distillation section (4).

Product quality: For gasoline production, typical properties of the Di-

mate are:
Specific gravity, @15°C 0.70
End point, °C 205 Utilities per ton of feed
70% vaporized, °C 80 Electric power, kWh 10.8
Rvp, bar 0.5 Steam, HP, t 0.14
RONC 96 Water, cooling, t 28.5
MONC 81 Catalyst + chemicals, US$ 9.3
RON blending value, avg. 103
Installation: Twenty-seven units have been built or are under construc-
Economics: For a plant charging 100,000 tpy of C3 cut (% propylene) tion.
and producing 71,000 tpy of Dimate gasoline:
Reference: “Olefin oligomerization with homogeneous catalysis,” 1999
Investment for a 2002 ISBL Gulf Coast erected cost,
Dewitt Petrochemical Conference, Houston.
excluding engineering fees, US $7 million
Licensor: Axens.
Oligomerization—polynaphtha
Application: To produce C6+ isoolefin fractions that can be used as high- ��
octane blending stocks for the gasoline pool and high-smoke-point
blending stocks for kerosine and jet fuel. The Polynaphtha and electopol
processes achieve high conversions of light olefinic fractions into higher
value gasoline and kerosine from propylene and mixed-butene fractions �
such as C3 and C4 cuts from cracking processes. �

Description: Propylene or mixed butenes (or both) are oligomerized

catalytically in a series of fixed-bed reactors (1). Conversion and selectiv-
ity are controlled by reactor temperature adjustment while the heat of
reaction is removed by intercooling (2). The reactor section effluent is
� ��
fractionated (3), producing raffinate, gasoline and kerosine. ��
The Selectopol process is a variant of the polynaphtha process where
the operating conditions are adjusted to convert selectively the isobu-
tene portion of an olefinic C4 fraction to high-octane, low-Rvp gasoline
blending stock. It provides a low-cost means of debottlenecking existing
alkylation units by converting all of the isobutene and a small percent-
age of the n-butenes, without additional isobutane. trogen or water contents in the feed warrant; however, the equipment
Polynaphtha and Selectopol processes have the following features: cost is low.
low investment, regenerable solid catalyst, no catalyst disposal problems,
long catalyst life, mild operating conditions, versatile product range,
Economics: Typical ISBL Gulf Coast investments for 5,000-bpd of FCC
C4 cut for polynaphtha (of maximum flexibility) and Selectopol (for maxi-
good-quality motor fuels and kerosine following a simple hydrogenation
mum gasoline) units are US$8.5 million and US$3.0 million, respectively.
step and the possibility of retrofitting old phosphoric acid units.
Respective utility costs are US$4.4 and US$1.8 per ton of feed while
The polygasoline RON and MON obtained from FCC C4 cuts are sig-
catalyst costs are US$0.2 per ton of feed for both processes.
nificantly higher than those of FCC gasoline and, in addition, are sulfur-
free. Hydrogenation improves the MON, whereas the RON remains high Installations: Seven Selectopol and polynaphtha units have been licensed (five
and close to that of C4 alkylate. in operation), with a cumulative operating experience exceeding 40 years.
Kerosine product characteristics such as oxidation stability, freezing
point and smoke point are excellent after hydrogenation of the polynaph- Licensor: Axens.
tha product. The kerosine is also sulfur-free and low in aromatics.
The Polynaphtha process has operating conditions very close to
those of phosphoric acid poly units. Therefore, an existing unit’s major
equipment items can be retained with only minor changes to piping
and instrumentation. Some pretreatment may be needed if sulfur, ni-
Paraxylene
Application: CrystPX is a modern crystallization technology to produce ��
��
high-purity paraxylene (PX). This process offers lower capital cost and ap-
��
plies a simpler process scheme when compared to other technologies. ��
��
��
CrystPX can be used in grassroots designs as a more economical alterna- ��
��
tive to adsorption processes, or applied in various revamp configurations ��
��
��
to improve product purity, increase capacity or lower operating costs. ��

Description: CrystPX uses reliable suspension crystallization as the meth-

��
od to produce PX from a mixture of C8 aromatics. The technology incor- ��
porates an optimized arrangement of equipment to conserve the cool-
��
ing energy and reduce recycle rates. A pusher-type centrifuge separates ��
��
PX crystals from the mother liquor, which is recycled to another stage, ��
��
or xylene isomerization unit. The number of stages required is set by
the feedstock composition and recovery required. The PX crystals are ��
washed with paraxylene product, avoiding the use of other components
that must subsequently be separated.
This process is economical to use with equilibrium xylene feedstock
(20 – 25% PX); or with more concentrated feeds, such as originating
from selective toluene conversion processes. In these cases, the process Process advantages include:
technology is even more economical to produce high-purity PX product. • High PX purity and recovery (99.8 + wt.% purity at up to 95%
This technology takes advantage of recent advances in crystallization recovery)
techniques and advancements in equipment to create this economically • Crystallization equipment is simple, easy to procure and opera-
attractive method for PX recovery and purification. tionally trouble free
The design uses only crystallizers and centrifuges in the primary op- • Compact design requires small plot size, and lowest capital invest-
eration. This simplicity of equipment promotes low maintenance costs, ment
easy incremental expansions and controlled flexibility. For the case with • Operation is flexible to meet market requirements for PX purity
concentrated feedstock, high-purity PX is produced in the front section • System is easily amenable to future requirement for incremental
of the process at warm temperatures, taking advantage of the high con- capacity increases
centration of PX present in the feed. At the back end of the process, ad- • Feed concentration of PX is used efficiently
ditional PX recovery is obtained through a series of crystallizers operated • Technology is flexible to process a range of feed concentrations
successively at colder temperatures. This scheme minimizes the need (20 – 95 wt.% PX) using a single or multistage system.
for recycling excessive amounts of filtrate, thus reducing total energy • Aromatics complex using CrystPX technology is cost competitive
requirements. with adsorption-based systems for PX recovery.
Continued 
Paraxylene, continued
Economics: Table 1 lists the benefits of the CrystPX process.
TABLE 1. Techno-economic comparison of CrystPX to conventional
technologies
Basis: 90% PX feed purity, 400,000 tpy of 99.9 wt% PX
CrystPX Other crystallization technologies
Investment cost, $MM 26.0 40.0
Paraxylene recovery, % 95 95
Electricity consumption, kWh/ ton PX 50 80
Operation mode Continuous Batch

Licensor: GTC Technology Inc.

Note: CrystPX is a proprietary process technology marketed and licensed
by GTC Technology Inc., in alliance with Lyondell Chemical Co.
Prereforming with feed
ultra purification ��
Application: Ultra-desulfurization and adiabatic-steam reforming of hy-
drocarbon feed from reﬁnery offgas or natural gas through LPG to naph-
tha feeds as a prereforming step in the route to hydrogen production.

Description: Sulfur components contained in the hydrocarbon feed are ��
converted to H 2 S in the HDS vessel and then fed to two desulfurization
vessels in series. Each vessel contains two catalyst types—the first for
bulk sulfur removal and the second for ultrapurification down to sulfur ��
levels of less than 1 ppb. ��
The two-desulfurization vessels are arranged in series in such a way ��
that either may be located in the lead position allowing online change ��
��
out of the catalysts. The novel interchanger between the two vessels ��
allows for the lead and lag vessels to work under different optimized
conditions for the duties that require two catalyst types. This arrange-
ment may be retrofitted to existing units.
Desulfurized feed is then fed to a fixed bed of nickel-based catalyst
that converts the hydrocarbon feed, in the presence of steam, to a prod-
uct stream containing only methane together with H 2, CO, CO 2 and Installation: CRG process technology covers 40 years of experience with
unreacted steam which is suitable for further processing in a conven- over 150 plants built and operated. Ongoing development of the cata-
tional fired reformer. The CRG prereformer enables capital cost savings lyst has lead to almost 50 such units since 1990.
in primary reforming due to reductions in the radiant box heat load. It
also allows high-activity gas-reforming catalyst to be used. The ability to Catalyst: The CRG catalyst is manufactured under license by Johnson
increase preheat temperatures and transfer radiant duty to the convec- Matthey Catalysts.
tion section of the primary reformer can minimize involuntary steam
production. Licensor: The process and CRG catalyst are licensed by Davy Process
Technology.
Operating conditions: The desulfurization section typically operates be-
tween 170 ° C and 420 ° C and the CRG prereformer will operate over a
wide range of temperatures from 250 ° C to 650 ° C and at pressures up
to 75 bara.
Pressure swing adsorption—rapid
cycle
Application: Hydrogen recovery from fuel gas and hydrogen contain-
ing offgas streams in refinery and chemical processes offers many po-
tential benefits, including process uplift, reduced H2 costs, avoided
H2 plant expansion and emissions reductions. It also requires a cost-
effective separation technology to be economical. Rapid-cycle pres-
sure swing adsorption (RCPSA) technology offers a more-compact,
less-expensive and more-energy-efficient solution for H2 recovery
compared to conventional PSA technology. This technology has been
jointly developed by ExxonMobil Research and Engineering Co. (EMRE)
and QuestAir Technologies. The resulting product—the “QuestAir H- ��
��
6200”— will have its first large-scale commercial application in 2007 ��
in an ExxonMobil Refinery.

Process description: Despite its widespread usage in industry, traditional

PSA processes have multiple inherent disadvantages. Slow cycle speeds
and relatively large adsorbent beads must be used to avoid fluidization tems made up of multiple vessels of beaded adsorbent, complex process
of the adsorbent bed, resulting in very large systems with high materials piping and multiple switching valves can be replaced with integrated
and vessel costs. In addition, networks of individual switching valves, modular skid-mounted QuestAir H-6200 plants that are up to 1/20th
with associated instrumentation, control systems and process piping, the size of a conventional PSA of equivalent capacity, and significantly
add complexity and cost to conventional PSA systems. lower cost. In addition, the RCPSA’s modular skid mounted design re-
RCPSA technology overcomes the inherent disadvantages of con- duces installation time and cost.
ventional PSA by using two proprietary technologies: structured adsor-
bents, which replace conventional beaded PSA adsorbents, and inte- Yields: Available upon request.
grated rotary valves, which replace solenoid-actuated valves. Structured
adsorbents provide mass transfer coefficients that are up to 100 times Utilities: Available upon request.
higher than beaded adsorbents used in conventional PSA; thus, signifi- References: “PSA technology hits the fast lane,” Chemical Processing,
cantly increasing the productivity of a unit volume of adsorbent bed. August 2003, Compressors and Industrial gases section.
The multi-port rotary valves are used for rapid and efficient switching
of gases between adsorbent beds, effectively capturing the increased Licensor: ExxonMobil Research and Engineering Co., and QuestAir
capacity of the structured adsorbent. Technologies Inc.
Multi-bed RCPSA systems can be efficiently packaged in an inte-
grated, modular rotating bed design. The net result is that large PSA sys-
Refinery offgas—purification
and olefins recovery ��
��
��
��
��
��
��
��
��
��
Application: Purification and recovery of olefins from FCC, RFCC and
DCC offgas.
��
��
Products: Hydrogen, methane, ethylene and LPG.
��
��
Description: Refinery offgas streams (ROG) from fluid catalytic cracker ��
(FCC) units, deep catalytic cracking (DCC) units, catalytic pyrolysis pro-
cess (CPP) units and coker units are normally used as fuel gas in re-
fineries. However, these streams contain significant amounts of olefins ��
��
��
��
(ethylene and propylene), which can be economically recovered. In fact, ��
��
��
many such streams can be recovered with project payout times of less
than one year. ��
Offgas-recovery units can be integrated with existing olefins units ��
��
or, if the flows are large enough, stand-alone units may be feasible.
��
Offgas-recovery units can be broken down into sections including feed ��
��
contaminant removal, ethylene recovery and propylene recovery. Feed ��
contaminants including acid gases, O2, NOx, arsine, mercury, ammonia, ��

nitrites, COS, acetylenes and water must be removed. It is critical that

the designer of the unit be experienced with feedstock pretreatment
since many of the trace components in the ROG streams can impact the taminants to remove in the ROG are the O2 and NOx, which are typically
ultimate product purity, catalyst performance and operational safety. removed by hydrogenation to H2O and NH3. Commercially available hy-
The ethylene recovery section can be a stand-alone unit where ei- drogenation catalysts cause significant loss of ethylene to ethane. BASF
ther dilute ethylene or polymer-grade ethylene (PGE) is produced; or a together with Shaw have developed a copper-based catalyst (R3-81),
unit where partially recovered streams suitable for integration into an which in sulfided form is capable of complete hydrogenation of the O2
ethylene plant recovery section are produced. and NOx.
Integration of treated ROG into an ethylene plant involves compres-
sion and treatment /removal of contaminants. If an ethylene product is Deoxo reactor. The Deoxo Reactor, in which the sulfided-copper catalyst
required (dilute ethylene or PGE), an additional section is needed that (R3-81) is used, serves a dual function. By removing the O2, NOx and
separates hydrogen, nitrogen and methane in a cold box, followed by a acetylene, it provides necessary purification of olefins but is also vital
demethanizer, a deethanizer; and for PGE and C2 splitter. toward the safety of the process. Without it, the formation of explosive
Removal of contaminants including acid gases, COS, RSH, NO2, NH3, deposits in and around the cold box can become an issue. An additional,
HCN, H2O, AsH3 and Hg is achieved by established processing methods,
depending on the concentrations in the ROG feed. The difficult con- Continued 
Refinery offgas—purification and olefins recovery,
continued
economic benefit of the Shaw Stone & Webster solution, comes from
the superior selectivity of the special catalyst— allowing deep removal of
O2 and NOx, with negligible ethylene loss.

Installation: Three units are currently operating. Several units are under
construction, and many units are under design.

Licensor: Shaw Stone & Webster Inc.

Resid catalytic cracking
Application: Selective conversion of gasoil and heavy residual feed-
stocks.
Products: High-octane gasoline, distillate and C3– C4 olefins.
Description: For residue cracking the process is known as R2R (reactor–2 re-
generators). Catalytic and selective cracking occurs in a short-contact-time
riser where oil feed is effectively dispersed and vaporized through a propri-
etary feed-injection system. Operation is carried out at a temperature con-
sistent with targeted yields. The riser temperature profile can be optimized
with the proprietary mixed temperature control (MTC) system.
Reaction products exit the riser-reactor through a high-efficiency,
close-coupled, proprietary riser termination device RSS (riser separator
stripper). Spent catalyst is pre-stripped followed by an advanced high-ef-
ficiency packed stripper prior to regeneration. The reaction product va-
por may be quenched to give the lowest dry gas and maximum gasoline
yield. Final recovery of catalyst particles occurs in cyclones before the
product vapor is transferred to the fractionation section.
Catalyst regeneration is carried out in two independent stages
equipped with proprietary air and catalyst distribution systems resulting is tailored to refiner’s needs and can include wide turndown flexibility.
in fully regenerated catalyst with minimum hydrothermal deactivation, Available options include power recovery, wasteheat recovery, flue-gas
plus superior metals tolerance relative to single-stage systems. These treatment and slurry filtration.
benefits are derived by operating the first-stage regenerator in a partial- Existing gas oil units can be easily retrofitted to this technology. Re-
burn mode, the second-stage regenerator in a full-combustion mode vamps incorporating proprietary feed injection and riser termination de-
and both regenerators in parallel with respect to air and flue gas flows. vices and vapor quench result in substantial improvements in capacity,
The resulting system is capable of processing feeds up to about 6 wt% yields and feedstock flexibility within the mechanical limits of the exist-
ConC without additional catalyst cooling means, with less air, lower cat- ing unit.
alyst deactivation and smaller regenerators than a single-stage regen-
erator design. Heat removal for heavier feedstocks (above 6 CCR) may
Installation: Shaw Stone & Webster and Axens have licensed 27 full-
technology R2R units and performed more than 150 revamp projects.
be accomplished by using a reliable dense-phase catalyst cooler, which
has been commercially proven in over 56 units. Reference: Meyers, R., Handbook of Petroleum Refining Process, Third Ed.
The converter vessels use a cold-wall design that results in mini-
mum capital investment and maximum mechanical reliability and safety. Licensor: Shaw Stone & Webster and Axens, IFP Group Technologies.
Reliable operation is ensured through the use of advanced fluidization
technology combined with a proprietary reaction system. Unit design
Slack wax deoiling
Application: Process to produce high-melting and low-oil containing ��
��
��
hard wax products for a wide range of applications.
��
��
Feeds: Different types of slack waxes from lube dewaxing units, includ- ��
��

ing macrocrystalline (parafﬁnic) and microcrystalline wax (from residual ��
oil). Oil contents typically range from 5–25 wt%. ��
��

��

Products: Wax products with an oil content of less than 0.5 wt%, except ��
��
��
��
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for the microcrystalline paraffins, which may have a somewhat higher ��
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oil content. The deoiled wax can be processed further to produce high- ��
quality, food-grade wax. ��
��
Description: Warm slack wax is dissolved in a mixture of solvents and ��
cooled by heat exchange with cold main filtrate. Cold wash filtrate ��
��
is added to the mixture, which is chilled to filtration temperature in
scraped-type coolers. Crystallized wax is separated from the solution in
a rotary drum filter (stage 1). The main filtrate is pumped to the soft-wax
solvent recovery section. Oil is removed from the wax cake in the filter
by thorough washing with chilled solvent.
The wax cake of the first filter stage consists mainly of hard wax and Utility requirements (slack wax feed containing 20 wt% oil, per metric
solvent but still contains some oil and soft wax. Therefore, it is blown off ton of feed):
the filter surface and is again mixed with solvent and repulped in an agi-
tated vessel. From there the slurry is fed to the filter stage 2 and the wax Steam, LP, kg 1,500
cake is washed again with oil-free solvent. The solvent containing hard Water, cooling, m3 120
wax is pumped to a solvent recovery system. The filtrate streams of filter Electricity, kWh 250
stage 2 are returned to the process, the main filtrate as initial dilution to
the crystallization section, and the wash filtrate as repulp solvent. Installation: Wax deoiling units have been added to existing solvent de-
The solvent recovery sections serve to separate solvent from the waxing units in several lube refineries. The most recent reference in-
hard wax respectively from the soft wax. These sections yield oil-free cludes the revamp of a dewaxing unit into two-stage wax deoiling; this
hard wax and soft wax (or foots oil). unit went onstream in 2005.

Licensor: Uhde GmbH.

SO2 removal, regenerative
Application: Regenerative scrubbing system to recover SO2 from flue gas
��
containing high SO2 levels such as gas from FCC regenerator or inciner- ��
��
ated SRU tail gas and other high SO2 applications. The LABSORB process ��
is a low pressure drop system and is able to operate under varying condi- ��
��
��
tions and not sensitive to variations in the upstream processes. ��
��
Products: The product from the LABSORB process is a concentrated SO2 ��
stream consisting of approximately 90% SO2 and 10% moisture. This ��
��
stream can be sent to the front of the SRU to be mixed with H2S and ��
��
form sulfur, or it can be concentrated for other marketable uses.
��
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Description: Hot dirty flue gas is cooled in a flue-gas cooler or waste- ��
heat recovery boiler prior to entering the systems. Steam produced can ��
be used in the LABSORB plant. The gas is then quenched to adiabatic ��
saturation (typically 50°C–75°C) in a quencher/pre-scrubber; it proceeds
to the absorption tower where the SO2 is removed from the gas. The
tower incorporates multiple internal and re-circulation stages to ensure
sufficient absorption.
A safe, chemically stable and regenerable buffer solution is con-
tacted with the SO2-rich gas for absorption. The rich solution is then scrubbing operating history, design and economics,” Sulfur 2000, San
piped to a LABSORB buffer regeneration section where the solution is Francisco, October 2000.
regenerated for re-use in the scrubber. Regeneration is achieved using Confuorto, Eagleson and Pedersen, “LABSORB, A regenerable wet
low-pressure steam and conventional equipment such as strippers, con- scrubbing process for controlling SO2 emissions,” Petrotech-2001, New
densers and heat exchangers. Delhi, January 2001.
Economics: This process is very attractive at higher SO2 concentrations Licensor: Belco Technologies Corp.
or when liquid or solid effluents are not allowed. The system’s buffer
loss is very low, contributing to a very low operating cost. Additionally,
when utilizing LABSORB as an SRU tail-gas treater, many components
normally associated with the SCOT process are not required; thus saving
considerable capital.

Installations: One SRU tail-gas system and two FCCU scrubbing systems.
Reference: Confuorto, Weaver and Pedersen, “LABSORB regenerative
Sour gas treatment
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Application: The WSA process (Wet gas Sulfuric Acid) treats all types of ��
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sulfur-containing gases such as amine and Rectisol regenerator offgas,
��
SWS gas and Claus plant tail gas in reﬁneries, gas treatment plants, ��
petrochemicals and coke chemicals plants. The WSA process can also be ��
��
applied for SOx removal and regeneration of spent sulfuric acid.
Sulfur, in any form, is efﬁciently recovered as concentrated commer- ��
��
cial-quality sulfuric acid. ��

Description: Feed gas is combusted and cooled to approximately

420°C in a waste heat boiler. The gas then enters the SO2 converter ��
containing one or several beds of SO2 oxidation catalyst to convert SO2 ��
��
to SO3. The gas is cooled in the gas cooler whereby SO3 hydrates to
��
H2SO4 (gas), which is finally condensed as concentrated sulfuric acid
(typically 98% w/w).
The WSA condenser is cooled by ambient air, and heated air may ��
be used as combustion air in the incinerator for increased thermal effi-
ciency. The heat released by incineration and SO2 oxidation is recovered
as steam. The process operates without removing water from the gas.
Therefore, the number of equipment items is minimized, and no liquid
waste is formed. Cleaned process gas leaving the WSA condenser is sent
to stack without further treatment.
The WSA process is characterized by:
• More than 99% recovery of sulfur as commercial-grade
sulfuric acid
• No generation of waste solids or wastewater
• No consumption of absorbents or auxiliary chemicals
• Efficient heat recovery ensuring economical operation
• Simple and fully automated operation adapting to variations in
feed gas flow and composition.

Installation: More than 50 units worldwide.

Licensor: Haldor Topsøe A/S.
Spent acid regeneration
Application: The WSA process (Wet gas Sulfuric Acid) treats spent sul-
furic acid from alkylation as well as other types of waste sulfuric acid ��
in the petrochemical and chemicals industry. Amine regenerator offgas
and /or refinery gas may be used as auxiliary fuel. The regenerated acid ��
��
��
will contain min. 98% H 2SO 4 and can be recycled directly to the alkyla- ��
��
tion process.
The WSA process is also applied for conversion of H 2S and removal ��
��
of SOx. ��
��
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Description: Spent acid is decomposed to SO2 and water vapor in an
incinerator using amine regenerator offgas or refinery gas as fuel. The
��
SO 2 containing flue gas is cooled in a waste-heat boiler and solid matter
originating from the acid feed is separated in an electrostatic precipita- ��
tor. By adding preheated air, the process gas temperature and oxygen
content are adjusted before the catalytic converter converting SO2 to
SO3. The gas is cooled in the gas cooler whereby SO3 is hydrated to
H 2SO4 (gas), which is finally condensed as 98% sulfuric acid.
The WSA condenser is cooled by ambient air. The heated air may be
used as combustion air in the burner for increased thermal efficiency. The
heat released by incineration and SO2 oxidation is recovered as steam. Installation: More than 50 WSA units worldwide, including 7 for spent
The process operates without removing water from the gas. There- acid regeneration.
fore, the number of equipment items is minimized and no liquid waste
is formed. This is especially important in spent acid regeneration where Licensor: Haldor Topsøe A/S.
SO 3 formed by the acid decomposition will otherwise be lost with the
wastewater.
The WSA process is characterized by:
• No generation of waste solids or wastewater
• No consumption of absorbents or auxiliary chemicals
• Efficient heat recovery ensuring economical operation
• Simple and fully automated operation adapting to variations in
feed flow and composition.
Spent lube oil re-refining
Application: The Revivoil process can be used to produce high yields of ��
premium quality lube bases from spent motor oils. Requiring neither
acid nor clay treatment steps, the process can eliminate environmental ��
��
and logistical problems of waste handling and disposal associated with ��
conventional re-refining schemes. ��
��
Description: Spent oil is distilled in an atmospheric flash distillation ��
column to remove water and gasoline and then sent to the Thermal ��
��
Deasphalting (TDA) vacuum column for recovery of gas oil overhead
and oil bases as side streams. The energy-efficient TDA column features
��
excellent performance with no plugging and no moving parts. Metals ��
and metalloids concentrate in the residue, which is sent to an optional
��
Selectopropane unit for brightstock and asphalt recovery. This scheme ��
��
is different from those for which the entire vacuum column feed goes
through a deasphalting step; Revivoil’s energy savings are significant,
and the overall lube oil base recovery is maximized. The results are sub-
stantial improvements in selectivity, quality and yields.
The final, but very important step for base oil quality is a specific
hydrofinishing process that reduces or removes remaining metals and
metalloids, Conradson Carbon, organic acids, and compounds con- Economics: The process can be installed stepwise or entirely. A simpler
taining chlorine, sulfur and nitrogen. Color, UV and thermal stability scheme consists of the atmospheric flash, TDA and hydrofinishing unit
are restored and polynuclear aromatics are reduced to values far below and enables 70 – 80% recovery of lube oil bases. The Selectopropane
the latest health thresholds. Viscosity index remains equal to or better unit can be added at a later stage, to bring the oil recovery to the 95%
than the original feed. For metal removal (> 96%) and refining-purifi- level on dry basis. Economics below show that for two plants of equal
cation duty, the multicomponent catalyst system is the industry’s best. capacity, payout times before taxes are two years in both cases.
Product quality: The oil bases are premium products; all lube oil base Investment: Basis 100,000 metric tpy, water-free, ISBL 2004 Gulf
specifications are met by Revivoil processing from Group 1 through Coast, million US$
Group 2 of the API basestocks definitions. Besides, a diesel can be ob-
tained, in compliance with the EURO 5 requirements (low sulfur). Configuration 1 (Atm. flash, TDA and Hydrofinishing units) 30
Configuration 2 (Same as above + Selectopropane unit) 35
Health & safety and environment: The high-pressure process is in line
with future European specifications concerning carcinogenic PNA com-
pounds in the final product at a level inferior to 5 wppm (less than 1
wt% PCA—IP346 method). Continued 
Spent oil lube re-refining, continued
Utilities: Basis one metric ton of water-free feedstock
Config. 1 Config. 2
Electrical power, kWh 45 55
Fuel, million kcal 0.62 0.72
Steam, LP, kg — 23.2
Steam, MP, kg 872 890
Water, cooling, m3 54 59

Installation: Ten units have been licensed using all or part of the Revivoil
Technology.

Licensor: Axens and Viscolube SpA.

Sulfur processing
Application: The D’GAASS Sulfur Degassing Process removes dissolved
H2S and H2Sx from produced liquid sulfur. Undegassed sulfur can create ��
odor problems and poses toxic and explosive hazards during the storage
and transport of liquid sulfur.

Description: Degasiﬁcation is accomplished in a pressurized vertical ves- ��

sel where undegassed sulfur is efficiently contacted with pressurized ��
��
��
process air (instrument or clean utility air). The contactor vessel may be ��
��
located at any convenient location. The undegassed sulfur is pumped to
the vessel and intimately contacted with air across special fixed vessel
internals. ��
��
Operation at elevated pressure and a controlled temperature ac-
��
celerates the oxidation of H2S and polysulfides (H2Sx) to sulfur. The de- ��
gassed sulfur can be sent to storage or directly to loading without ad-
ditional pumping. Operation at elevated pressure allows the overhead
vapor stream to be routed to the traditional incinerator location, or to
the SRU main burner or TGTU line burner—thus eliminating the degas-
sing unit as an SO2 emission source.

Economics: D’GAASS achieves 10 ppmw combined H2S/H2Sx in product Reference: US Patent 5,632,967.
sulfur without using catalyst. Elevated pressure results in the following Nasato, E. and T. A. Allison, “Sulfur degasification—The D’GAASS
benefits: low capital investment, very small footprint, low operating cost process,” Laurance Reid Gas Conditioning Conference, Norman, Okla-
and low air requirement. Operation is simple, requiring minimal opera- homa, March, 1998.
tor and maintenance time. No chemicals, catalysts, etc., are required. Fenderson, S., “Continued development of the D’GAASS sulfur de-
gasification process,” Brimstone Sulfur Recovery Symposium, Canmore,
Installations: Twenty-one D’GAASS units in operation. Twenty-six ad-
Alberta, May 2001.
ditional trains in engineering and construction phase with total capacity
over 25,000 long ton per day (LTPD). Licensor: Goar, Allison & Associates, Inc.
Sulfur recovery
Application: The COPE Oxygen Enrichment Process allows existing Claus ��
��
sulfur recovery/tail gas cleanup units to increase capacity and recov- ��
ery, can provide redundant sulfur processing capacity, and can improve ��
��
combustion performance of units processing lean acid gas. � ��
��
Description: The sulfur processing capacity of typical Claus sulfur re- ��
��
��
covery units can be increased to more than 200% of the base capacity ��
��
through partial to complete replacement of combustion air with pure ��
oxygen (O2). SRU capacity is typically limited by hydraulic pressure drop. ��
��
As O2 replaces combustion air, the quantity of inert nitrogen is reduced ��
allowing additional acid gas to be processed. � ��
��
The process can be implemented in two stages. As the O2 enrich-
��
ment level increases, the combustion temperature (1) increases. COPE ��
Phase I, which does not use a recycle stream, can often achieve 50% ��
capacity increase through O2 enrichment to the maximum reaction fur-
nace refractory temperature limit of 2,700ºF – 2,800ºF. Higher O2 enrich-
ment levels are possible with COPE Phase II which uses an internal pro-
cess recycle stream to moderate the combustion temperature allowing
enrichment up to 100% O2. Operating costs are a function of O2 cost, reduced incinerator fuel, and
Flow through the remainder of the SRU (2, 3, and 4) and the tail reduced operating and maintenance labor costs.
gas cleanup unit is greatly reduced. Ammonia and hydrocarbon acid
gas impurity destruction and thermal stage conversion are improved at Installations: Twenty-nine COPE trains at 17 locations.
the higher O2 enriched combustion temperatures. Overall SRU sulfur Reference: US Patents 4,552,747 and 6,508,998.
recovery is typically increased by 0.5% to 1%. A single proprietary COPE Sala, L., W. P. Ferrell and P. Morris, “The COPE process—Increase
burner handles acid gas, recycle gas, air and O2. sulfur recovery capacity to meet changing needs,” European Fuels Week
Operating conditions: Combustion pressure is from 6 psig to 12 psig; Conference, Giardini Naxos, Taormina, Italy, April 2000.
combustion temperature is up to 2,800ºF. Oxygen concentration is from Nasato, E. and T. A. Allison, “COPE Ejector—Proven technology,”
21% to 100%. Sulphur 2002, Vienna, Austria, October 2002.
Licensor: Goar, Allison & Associates, Inc., and Air Products and Chemi-
Economics: Expanded SRU and tail gas unit retrofit sulfur processing
cals, Inc.
capacity at capital cost of 15% – 25% of new plant cost. New plant sav-
ings of up to 25%, and redundant capacity at 15% of base capital cost.
Thermal gasoil
Application: The Shell Thermal Gasoil process is a combined residue and
waxy distillate conversion unit. The process is an attractive low-cost con- ��
version option for hydroskimming refineries in gasoil-driven markets or ��
for complex refineries with constrained waxy distillate conversion capac-
ity. The typical feedstock is atmospheric residue, which eliminates the � ��
��
need for an upstream vacuum flasher. This process features Shell Soaker
Visbreaking technology for residue conversion and an integrated recycle
heater system for the conversion of waxy distillate.
��
� � ��
Description: The preheated atmospheric (or vacuum) residue is charged
to the visbreaker heater (1) and from there to the soaker (2). The con- ��
�
version takes place in both the heater and soaker and is controlled by
the operating temperature and pressure. The soaker effluent is routed ��
to a cyclone (3). The cyclone overheads are charged to an atmospheric ��
fractionator (4) to produce the desired products including a light waxy
distillate. The cyclone and fractionator bottoms are routed to a vacuum
flasher (6), where waxy distillate is recovered. The combined waxy distil-
lates are fully converted in the distillate heater (5) at elevated pressure. Utilities, typical consumption/production for a 25,000-bpd unit,
Yields: Depend on feed type and product specifications. dependent on configuration and a site’s marginal economic values for
steam and fuel:
Feed atmospheric residue Middle East
Fuel as fuel oil equivalent, bpd 675
Viscosity, cSt @ 100°C 31
Power, MW 1.7
Products, % wt. Net steam production (18 bar), tpd 370
Gas 6.4
Gasoline, ECP 165°C 12.9 Installation: To date, 12 Shell Thermal Gasoil units have been built. Post
Gasoil, ECP 350°C 38.6 startup services and technical services for existing units are available
Residue, ECP 520°C+ 42.1 from Shell Global Solutions..
Economics: The typical investment for a 25,000-bpd unit will be about Licensor: Shell Global Solutions International B.V., and ABB Lummus
$2,400 to $3,000/bbl installed, excluding treating facilities. (Basis: West- Global B.V.
ern Euope, 2004.)
Treating—jet fuel/kerosine
Application: NAPFINING / MERICAT II / AQUAFINING systems eliminate ��
��
naphthenic acids and mercaptans from kerosine to meet acid number ��
and mercaptan jet fuel specifications. Caustic, air and catalyst are used ��
along with FIBER-FILM Contactor technology and an upflow catalyst im- ��
pregnated carbon bed saturated with caustic.

Description: In the NAPFINING system, the caustic phase ﬂows along

the fibers of the FIBER-FILM Contactor as it preferentially wets the fi- ��
��
bers. The kerosine phase simultaneously flows through the caustic-wet-
ted fibers where naphthenic acids react with the caustic phase to form
sodium naphthenate. The two phases disengage and the naphthenic
acid-free kerosine flows to the MERICAT II where the mercaptans react
with caustic, air, and catalyst in the FIBER-FILM Contactor to form disul- ��
fides. The two phases disengage again and the kerosine flows upwards ��
��
through a catalyst impregnated carbon bed where the remaining heavy
mercaptans are converted to disulfides.
An AQUAFINING system is then used to water wash the kerosine
downstream of the MERICAT II vessel to remove sodium. Salt driers and
clay filters are used downstream of the water wash to remove water,
surfactants and particulates to ensure a completely clean product.

Competitive advantages: FIBER-FILM Contactor technology requires

smaller processing vessels thus saving valuable plant space and reduc-
ing capital expenditure. Onstream factor is 100% whereas electrostatic
precipitators and downﬂow ﬁxed-bed reactors onstream factors are un-
predictable.

Installation: One hundred seventy-ﬁve installations worldwide.

References: Hydrocarbon Technology International, 1993.
Petroleum Technology Quarterly, Winter 1996/97.

Licensor: Merichem Chemicals & Reﬁnery Services LLC.

Treating—gases ��
��
Application: AMINEX and THIOLEX systems extract H2S from gases with ��
��
amine or caustic solution using FIBER-FILM Contactor technology.

Description: In an AMINEX system, the amine phase ﬂows along the ﬁ-

bers of the FIBER-FILM Contactor as it preferentially wets the fibers. The
gas phase flows through the Contactor parallel to the amine-wetted
fibers as the H2S is extracted into the amine. The two phases disengage
in the separator vessel with the rich amine flowing to the amine regen-
eration unit and the treated gas flowing to its final use.
Similarly, a THIOLEX system employs the same process utilizing caus-
tic to preferentially wet the fibers as the H2S is extracted into the caustic
��
phase. The rich caustic flows to sulfidic caustic storage and the treated
��
gas flows to its final use. ��
��
Competitive advantages: FIBER-FILM Contactor technology requires
smaller processing vessels thus saving valuable plant space and reducing
capital expenditures.

Installation: Four installations worldwide in THIOLEX service.

Reference: Hydrocarbon Processing, Vol. 63, No. 4, April 1984, P. 87.
Licensor: Merichem Chemicals & Refinery Services LLC.
Treating— gasoline and LPG ��
Application: THIOLEX/REGEN systems extract H2S and mercaptans from ��
gases and light liquid hydrocarbon streams, including gasolines, with ��
caustic using FIBER-FILM Contactor technology. It can also be used to
hydrolyze and remove COS from LPG and propane. ��
��
��
Description: In a THIOLEX system, the caustic phase flows along the
fibers of the FIBER-FILM Contactor as it preferentially wets the fibers.
Hydrocarbon flows through the caustic-wetted fibers where the H2S and
mercaptans are extracted into the caustic phase. The two phases disen-
��
gage and the caustic flows to the REGEN where the caustic is regener- ��
��
ated using heat, air and catalyst. The disulfide oil formed in this reaction ��
may be removed via gravity separation, FIBER-FILM solvent washing or
��
a combination of the two. The regenerated caustic flows back to the ��
THIOLEX system for continued re-use.
��
COS is removed from LPG or propane by either employing AMINEX
technology using an amine solution or THIOLEX technology using an
MEA/caustic solution to hydrolyze the COS to H2S and CO2 which are
easily removed by amine or caustic.

Competitive advantages: FIBER-FILM Contactor technology requires

smaller processing vessels thus saving valuable plant space and reducing
capital expenditures.

Installation: At present, 350 installations worldwide.

References: Oil & Gas Journal, August 12, 1985, p. 78.
Hydrocarbon Engineering, February 2000.

Licensor: Merichem Chemicals & Reﬁnery Services LLC.

Treating—gasoline desulfurization,
ultra deep ��
��
��
Application: EXOMER extracts recombinant mercaptan sulfur from se- ��
lectively hydrotreated FCC gasoline streams with a proprietary treating ��
solution. FIBER-FILM Contactor technology is used for mass transfer ef- ��
ficiency to obtain a maximum reduction in total sulfur content. EXOMER ��
is jointly developed with ExxonMobil Research & Engineering Co. ��
��
��
Description: In an EXOMER system, the lean treating solution phase
flows along the fibers of the FIBER-FILM Contactor along with the hy- ��
drocarbon phase, allowing the recombinant mercaptans to be extracted ��
��
into the treating solution in a non-dispersive manner. The two phases ��
disengage in the separator vessel with the treated hydrocarbon flowing ��
��
to storage. ��
The separated rich treating solution phase is sent to the regenera- ��
tion unit where sulfur-bearing components are removed. The removed
sulfur is sent to another refinery unit for further processing. The regen-
erated lean treating solution is returned to the EXOMER extraction step
for further use.

Economics: EXOMER allows reﬁners to meet stricter sulfur speciﬁca-

tions while preserving octane by allowing the hydrotreater severity to be
reduced. The capital expenditure for a grass roots EXOMER is 35 – 50%
of the cost of incremental hydrotreating capacity. Operating costs per
barrel are about 60 –70% less than hydrotreating.

Installation: Three installations worldwide.

Reference: Hydrocarbon Processing, February 2002, p. 45.
Licensor: Merichem Chemicals & Reﬁnery Services LLC.
Treating— gasoline sweetening
��
Application: MERICAT systems oxidize mercaptans to disulﬁdes by react-
ing mercaptans with air and caustic in the presence of catalyst using
FIBER-FILM Contactor technology.

Description: In a MERICAT system, the caustic phase flows along the fi- ��
bers of the FIBER-FILM Contactor as it preferentially wets the fibers. Prior
to entering the FIBER-FILM Contactor the gasoline phase mixes with air
through a proprietary air sparger. The gasoline then flows through the
caustic-wetted fibers in the Contactor where the mercaptans are ex- ��
��
tracted and converted to disulfides in the caustic phase. The disulfides
are immediately absorbed back into the gasoline phase. The two phases
disengage and the caustic is recycled back to the FIBER-FILM Contactor
until spent. ��

Competitive advantages: FIBER-FILM Contactor technology uses smaller

processing vessels while guaranteeing the sodium content of the prod-
uct. This saves valuable plant space and reduces capital expenditure.

Installation: At present, 135 installations worldwide.

Licensor: Merichem Chemicals & Reﬁnery Services LLC.
Treating—kerosine and heavy naphtha
sweetening ��

Application: MERICAT II oxidizes mercaptan sulfur to disulﬁdes to reduce

product odor. The streams treated are jet fuel, kerosine, natural gasoline ��
and selectively hydrotreated FCC gasolines.
��
Description: A MERICAT II system consists of two treaters. The FIBER-
FILM Contactor section removes hydrogen sulfide and naphthenic acids
while converting some mercaptans to disulfides with air, oxidation cata-
lyst and caustic solution. The partially treated hydrocarbon exits the FI-
BER-FILM Contactor and passes upflow through a catalyst-impregnated
carbon bed saturated with caustic to convert the remaining high-boiling ��
��
mercaptans to disulfides.

Competitive advantages:
• Minimal caustic and catalyst consumption
• Operating simplicity
• Minimal capital investment
• Recausticizing of the carbon bed without interruption of treating.
The FIBER-FILM section keeps organic acids from entering the carbon
bed. This conserves caustic and avoids fouling of the bed with sodium
naphthenate soaps. Competitive downﬂow reactors need frequent car-
bon bed hot water washings to remove these soaps whereas MERICAT
II does not require hot water washes.
The MERICAT II onstream factor is 100% while competitive systems
requiring periodic cleaning have unpredictable onstream factors.

Installation: Thirty-four installations worldwide.

Reference: Hydrocarbon Technology International, 1993.
Licensor: Merichem Chemicals & Reﬁnery Services LLC.
Treating—phenolic caustic ��

Application: ECOMERICAT removes phenols from phenolic caustics by ��

neutralization in conjunction with solvent washing using a FIBER-FILM ��

Contactor.

Description: An ECOMERICAT system contacts the spent caustic with a

slipstream of sweetened gasoline containing CO2 whereby neutralizing
��
the spent caustic, springing the phenols, absorbing the phenols into
the sweetened gasoline yielding neutral brine with minimal phenolic ��

content.
��
Competitive advantages: ��
��
• Minimizes spent caustic disposal cost ��
• Reduces the phenol content of spent caustic and increases the ��
��
phenol content of sweetened gasoline thus adding value
• Operates over a wide pH range ��
��
• Simple to operate
• No corrosion problems due to the buffering effect of CO2.

Installation: One installation worldwide.

Licensor: Merichem Chemicals & Reﬁnery Services LLC.
Treating—Pressure swing adsorption
Application: Pressure swing adsorption (PSA) process selectively adsorb-
ing impurities from product streams. The impurities are adsorbed in a ��
ﬁxed-bed adsorber at high pressure and desorbed by “swinging” the ��
adsorber from the feed to the tail gas pressure and by using a high-pu-
rity purge. The desired component is not adsorbed and is recovered at � � � � �
high purity.

Description: A PSA system operates as a batch process. However, mul-

tiple adsorbers operating in a staggered sequence are used to produce
constant feed, product and tail gas flows.
Step 1: Adsorption. The feed gas enters an adsorber at a high pres- ��
sure, impurities are adsorbed and high-purity product is produced. Flow
��
is normally in the upwardly direction. When an adsorber has reached ��
its adsorption capacity, it is taken offline, and the feed automatically
switched to a fresh adsorber.
Step 2: Co-current depressurization. To recover the product trapped
in the adsorbent void spaces, the adsorber is co-currently (in the direc-
tion of feed flow) depressurized. The product gas withdrawn is used
internally to repressurize and purge other adsorbers.
Step 3: Counter-current depressurization. At the end of the co-cur- UOP’s polybed PSA system offers:
rent depressurization step, the adsorbent is partially regenerated by • High reliability (greater than 99.8% onstream time)
counter-currently depressurizing the adsorber to the tail-gas pressure, • Minimal manpower requirements due to automatic operation
and thereby rejecting the impurities. • Reduced equipment costs and enhanced performance based on
Step 4: Purge. The adsorbent is purged with a high-purity stream high performance
(taken from another adsorber on the cocurrent depressurization step) at • Adsorbents and advanced PSA cycles
a constant pressure to further regenerate the bed. • Lower operating and equipment costs for downstream process
Step 5: Repressurization. The repressurization gas is provided from units
the co-current depressurization step and a slipstream from the product. • Flexibility to process more than one feedstock
When the adsorber has reached the adsorption pressure, the cycle has • Minimal feed pretreatment and utility requirements
been completed. The vessel is ready for the next adsorption cycle. • Adsorbents last for the life of the mechanical equipment (more
than 30 years).

Continued 
Treating—pressure swing adsorption, continued
Installation: Since commercialization in 1966, UOP has provided over
700 PSA systems in more than 60 countries in the refining, petrochemi-
cal, polymer, steel and power-generation industries. The Polybed PSA
system has demonstrated exceptional economic value in many appli-
cations, such as hydrogen recovery from refinery off-gases, monomer
recovery monomers in polyolefin plants, hydrogen extraction from gas-
ification syngas, helium purification for industrial gas use, adjustment
of synthesis gas for ammonia production, methane purification for pet-
rochemicals production, and H2/CO ratio adjustment for syngas used in
the manufacture of oxo-alcohols. Feed conditions typically range from
(7–70 kg/cm2g) (100 to 1,000 psig), with concentrations of the desired
component from 30 – 98+ mol %. System capacities range from less than
1 to more than 350 MMscfd (less than 1,100 to more than 390,000
Nm3/h).

Licensor: UOP LLC.

Treating—propane
Application: AMINEX extracts H2S and COS from propane with an amine
solution using FIBER-FILM Contactor technology. ��

Description: In an AMINEX system, the amine phase ﬂows along the ﬁ-

bers of the FIBER-FILM Contactor as it preferentially wets the fibers. The
propane phase flows through the amine-wetted fibers as the H2S and
COS are extracted into the amine phase. The two phases disengage in
the separator vessel with the rich amine flowing to the amine regenera-
tion unit and the treated propane flowing to storage.

Competitive advantages: FIBER-FILM Contactor technology requires

smaller processing vessels thus saving valuable plant space and reducing
capital expenditure. ��
��

Installation: Twenty installations worldwide.

Reference: Hydrocarbon Processing, Vol. 63, No. 4, April 1984, p. 87.
Licensor: Merichem Chemicals & Reﬁnery Services LLC.
Treating—reformer products
Application: CHLOREX removes inorganic chloride compounds from liq-
��
uid and gas reformer products using a FIBER-FILM Contactor and an
alkaline water treating solution. ��

Description: The CHLOREX system uses an alkaline water solution to

extract chloride impurities contained in the reformate stabilizer feed or
the stabilizer overhead product. CHLOREX can also be used to remove
��
chlorides from reformer offgas. Fresh caustic and fresh process water are
added to the system to maintain the proper pH of the recycle solution.

Competitive advantages: CHLOREX produces an easily handled waste ��

when compared to disposal of sacriﬁcial solid bed absorbents.
��
Installation: Four installations worldwide. ��

Licensor: Merichem Chemicals & Reﬁnery Services LLC.

Treating— spent caustic deep
neutralization ��
��
Application: MERICON systems neutralize spent caustics containing sul- ��
fides, mercaptans, naphthenic acids, and phenols. ��
��
Description: A MERICON system neutralizes spent caustic with acid to
a low pH. The sprung acid oils are separated from the acidic brine. The
resulting acid gases (H2S and mercaptans) flow to a sulfur plant. The ��
sprung acid oils are returned to the refinery for processing. The acid- ��
ic brine is further stripped with fuel gas to remove traces of H2S and
��
mercaptans. Finally, the acidic brine is mixed with caustic to return it to ��
a neutral pH for final disposal.
��
Competitive advantages:
��
• Minimal operator attention and 100 % onstream factor between
turnarounds
• Minimal capital investment
• Maximum COD reduction
• Non-odorous neutralized brine product
• Recovery of valuable hydrocarbons.

Installation: Seventeen installations worldwide.

Reference: Petroleum Technology Quarterly, Spring 2001, p. 55.
Licensor: Merichem Chemicals & Reﬁnery Services LLC.
Vacuum distillation
Application: Process to produce vacuum distillates that are suitable for ��
lubricating oil production by downstream units, and as feedstocks to ��
FCC and hydrocracker units.

Feed: Atmospheric bottoms from crude oils (atmospheric residue) or ��

hydrocracker bottoms. ��

Product: Vacuum distillates of precisely defined viscosities and flash points ��
��
(for lube production) and low metals content (for FCC and hydrocracker ��
units) as well as vacuum residue with specified softening point, penetra- ��
tion and flash point. ��

Description: Feed is preheated in a heat-exchanger train and fed to the ��

fired heater. The heater outlet temperature is controlled to produce the
��
required quality of vacuum distillates and residue. Structured packings
are typically used as tower internals to achieve low flashzone pressure
and, hence, to maximize distillate yields. Circulating reflux streams en-
able maximum heat recovery and reduced column diameter.
A wash section immediately above the flash zone ensures that the
metals content in the lowest side draw is minimized. Heavy distillate Installation: Numerous installations using the Uhde (Edeleanu) propri-
from the wash trays is recycled to the heater inlet or withdrawn as met- etary technology are in operation worldwide. The most recent reference
als cut. is a 86,000-bpd unit for a German refinery, which was commissioned
When processing naphthenic residues, a neutralization section may in 2004; the unit produces vacuum distillates as feedstock for FCC and
be added to the fractionator. hydrocracker units.
Utility requirements (typical, North Sea Crude), units per m³ of feed: Licensor: Uhde GmbH.
Electricity, kWh 5
Steam, MP, kg 15
Steam production, LP, kg 60
Fuel oil, kg 7
Water, cooling, m³ 3
Visbreaking
Application: Manufacture incremental gas and distillate products and simul- ��
taneously reduce fuel oil viscosity and pour point. Also, reduce the amount
of cutter stock required to dilute the resid to meet the fuel oil specifications.
Foster Wheeler/UOP offer both “coil” and “soaker” type visbreaking pro-
cesses. The following information pertains to the “coil” process.
��
Products: Gas, naphtha, gas oil, visbroken resid (tar). ��
��
Description: In a “coil” type operation, charge is fed to the visbreaker ��
heater (1) where it is heated to a high temperature, causing partial va-
porization and mild cracking. The heater outlet stream is quenched with ��
gas oil or fractionator bottoms to stop the cracking reaction. The va-
por-liquid mixture enters the fractionator (2) to be separated into gas, ��
naphtha, gas oil and visbroken resid (tar). The tar may also be vacuum
flashed for recovery of visbroken vacuum gas oil.

Operating conditions: Typical ranges are:

Heater outlet temperature, ºF 850 – 910
Quenched temperature, ºF 710 – 800
An increase in heater outlet temperature will result in an increase in Economics:
overall severity, further viscosity reduction and an increase in conversion. Investment (basis: 40,000 – 10,000 bpsd, 2Q 2005, US Gulf),
$ per bpsd 800 – 1,800
Yields: Utilities, typical per bbl feed:
Feed, source Light Arabian Light Arabian Fuel, MMBtu 0.1195
Type Atm. Resid Vac. Resid Power, kW/bpsd 0.0358
Gravity, ºAPI 15.9 7.1 Steam, MP, lb 6.4
Sulfur, wt% 3.0 4.0 Water, cooling, gal 71.0
Concarbon, wt% 8.5 20.3
Installation: Over 50 units worldwide.
Viscosity, CKS @130ºF 150 30,000
CKS @ 210ºF 25 900 Reference: Handbook of Petroleum Refining Processes, Third Ed., Mc-
Products, wt% Graw-Hill, 2003, pp. 12.91 – 12.105.
Gas 3.1 2.4
Licensor: Foster Wheeler/UOP LLC.
Naphtha (C5 – 330 ºF) 7.9 6.0
Gas oil 14.5 15.5
Visbroken resid 74.5 (600ºF+) 76.1 (662ºF+)
Visbreaking
��
Application: The Shell Soaker Visbreaking process is most suitable to re-
duce the viscosity of vacuum (and atmospheric) residues in (semi) complex
refineries. The products are primarily distillates and stable fuel oil. The to-
tal fuel oil production is reduced by decreasing the quantity of cutter stock � ��
required. Optionally, a Shell vacuum flasher may be installed to recover
additional gas oil and waxy distillates as cat cracker or hydrocracker feed ��
� ��
from the cracked residue. The Shell Soaker Visbreaking technology has
��
also proven to be a very cost-effective revamp option for existing units. ��
��
Description: The preheated vacuum residue is charged to the visbreaker �
��
heater (1) and from there to the soaker (2). The conversion takes place �
in both the heater and the soaker. The operating temperature and pres-
sure are controlled such as to reach the desired conversion level and/ ��
or unit capacity. The cracked feed is then charged to an atmospheric
fractionator (3) to produce the desired products like gas, LPG, naph-
tha, kerosine, gas oil, waxy distillates and cracked residue. If a vacuum
flasher is installed, additional gas oil and waxy distillates are recovered bpd unit, dependent on configuration and a site’s marginal economic
from the cracked residue. values for steam and fuel:
Yields: Vary with feed type and product specifications. Fuel as fuel oil equivalent, bpd 400
Power, MW 1.2
Feed, vacuum residue Middle East Net steam production (18 bar), tpd 370
Viscosity, cSt @100°C 615
Products, wt% Installation: More than 70 Shell Soaker Visbreakers have been built. Post
Gas 2.2 startup services and technical services for existing units are available from
Gasoline, 165°C EP 4.8 Shell Global Solutions.
Gas oil, 350°C EP 13.6
Licensor: Shell Global Solutions International B.V. and ABB Lummus
Waxy distillate, 520°C EP 23.4
Global B.V.
Residue, 520°C+ 56

Economics: The typical investment for a 25,000-bpd unit will be about

$1,200 to $1,500/bbl installed, excluding treating facilities. (Basis: West-
ern Europe, 2004.)
Utilities, typical consumption consumption/production for a 25,000-
Wax hydrotreating
Application: Hydrogen finishing technology has largely replaced clay
treatment of low-oil-content waxes to produce food- and medicinal- ��
grade product specifications (color, UV absorbency and sulfur) in new
units. Advantages include lower operating costs, elimination of environ- �
mental concerns regarding clay disposal and regeneration, and higher �
net wax product yields.
Bechtel has been offering for license the Wax Hy-Finishing process. ��
Bechtel now is marketing a line of modular, standard hydrogen finish- ��
��

ing units for wax treatment. Standard sizes are 500, 1,000, 2,000 and �
3,000-bpsd feedrate.
��
The core of the unit is standardized; however, individual modules are
modiﬁed as needed for speciﬁc client needs. This unit will be fabricated
to industry standards in a shop environment and delivered to the plant ��
site as an essentially complete unit. Cost and schedule reductions of at
least 20% over conventional stick-built units are expected. The standard
licensor’s process guarantees and contractor’s performance guarantees
(hydraulic and mechanical) come with the modules.

Description: Hard-wax feed is mixed with hydrogen (recycle plus make-

up), preheated, and charged to a fixed-bed hydrotreating reactor (1). Utilitiies, typical per bbl feed:
The reactor effluent is cooled in exchange with the mixed feed-hydrogen Fuel, 103 Btu (absorbed) 30
stream. Gas-liquid separation of the effluent occurs first in the hot sepa- Electricity, kWh 5
rator (2) then in the cold separator (3). The hydrocarbon liquid stream Steam, lb 25
from each of the two separators is sent to the product stripper (4) to Water, cooling (25°F rise), gal 300
remove the remaining gas and unstabilized distillate from the wax prod-
uct, and the product is dried in a vacuum flash (5). Gas from the cold Licensor: Bechtel Corp.
separator is either compressed and recycled to the reactor or purged
from the unit if the design is for once-through hydrogen.

Economics:
Investment (Basis 2,000-bpsd feedrate capacity,
2006 US Gulf Coast), $/bpsd 7,300
Wet gas scrubbing (WGS) ��

Application: To reduce ﬂuid catalytic cracking unit (FCCU) particulate ��

(catalyst) and sulfur oxides (SOx) emissions to achieve compliance with ��
environmental regulations—generally NSPS (New Source Performance
Standards) and consent decrees (in the US). The technology can also be
��
adapted for other refinery applications, e.g., coke calciners when flue ��
��
��
gases must be treated to reduce SOx and particulate emissions. ��
��
Description: The WGS process takes dirty gas from FCCUs and simul- ��
taneously removes particulate matter and SOx via direct contact with a ��
buffered liquid. Particulate removal is accomplished via inertial impaction ��
��
of the particulate with the scrubbing liquid. SO2 removal is accomplished ��
��
via absorption into and reaction with the buffered scrubbing liquid. SO3
removal is accomplished via a combination of nucleate condensation,
absorption and inertial impaction. All of this can be accomplished with ��
low- (3-in. of water column) or no pressure drop. This is important when ��
aged heat recovery systems are involved. ��
��
The liquid purge from the WGS system is further treated in either the ��
refinery wastewater system or in a segregated system, either of which
will remove solids for landfill disposal and will reduce the chemical oxy-
gen demand (COD) of the stream to meet local discharge requirements. requirements of less than 10 vppmd. In addition, the WGS has recorded
Operation of WGS systems demonstrates: run lengths in excess of 10 years without affecting FCCU throughput.
• Flexible/”forgiving” performance under a wide range of FCCU op- Installation: Nineteen operating plants have over 400 years of operating
erations/upsets experience. Four additional units are in various stages of engineering.
• Service factors equal to or better than FCCUs with runs exceeding
10 years Reference: 1991 AIChE Spring National Meeting, Paper No. 62c.
• Low/zero pressure drop 1990 National Petroleum Refiners Association Annual Meeting, Pa-
• Ability to meet stringent emission regulations, e.g., consent de- per No. AM-90-45, March 1990.
cree requirements. 1996 National Petroleum Refiners Association Annual Meeting, Pa-
per No. AM-96-47, March 1996.
Performance: All WGS facilities are in compliance with their permitted val-
ues. Compliance has been achieved for the current consent decree particu- Technology owner: ExxonMobil Research and Engineering Co.
late limits that can require emission of less than ½-lb of particulates /1,000 Licensor: Hamon Research-Cottrell, GN-Hamon, LLC.
lb of coke burned. They are also in compliance with consent decree SO2
requirements of 25 vppmd @ 0% O2 and SO3 consent decree emission
Wet scrubbing system, EDV
Application: EDV Technology is a low pressure drop scrubbing system, ��
to scrub particulate matter (including PM2.5), SO2 and SO3 from flue ��
gases. It is especially well suited where the application requires high reli-
ability, flexibility and the ability to operate for 4 – 7 years continuously ��
without maintenance shutdowns. The EDV technology is highly suited ��
��
for FCCU regenerator flue-gas applications. ��
��
Products: The effluents from the process will vary based on the re-
agent selected for use with the scrubber. In the case where a sodium- ��
based reagent is used, the product will be a solution of sodium salts. ��
��
Similarly, a magnesium-based reagent will result in magnesium salts. ��
��
A lime/limestone-based system will produce a gypsum waste. The EDV
technology can also be used with the LABSORB buffer thus making ��
the system regenerative. The product, in that case, would be a usable ��

condensed SO 2 stream.

Description: The flue gas enters the spray tower through the quench
section where it is immediately quenched to saturation temperature.
It proceeds to the absorber section for particulate and SO2 reduction. Economics: The EDV wet scrubbing system has been extremely suc-
The spray tower is an open tower with multiple levels of BELCO-G- cessful in the incineration and refining industries due to the very high
Nozzles. These nonplugging and abrasion-resistant nozzles remove scrubbing capabilities, very reliable operation and reasonable price.
particulates by impacting on the water/reagent curtains. At the same Installation: More than 200 applications worldwide on various pro-
time, these curtains also reduce SO 2 and SO 3 emissions. The BELCO- cesses including 43 FCCU applications, 5 heater applications, 1 SRU
G-Nozzles are designed not to produce mist; thus a conventional mist tailgas unit and 1 fluidized coker application to date.
eliminator is not required.
Upon leaving the absorber section, the saturated gases are direct- Reference: Confuorto and Weaver, “Flue gas scrubbing of FCCU re-
ed to the EDV filtering modules to remove the fine particulates and ad- generator flue gas—performance, reliability, and flexibility—a case his-
ditional SO3. The filtering module is designed to cause condensation of tory,” Hydrocarbon Engineering, 1999.
the saturated gas onto the fine particles and onto the acid mist, thus Eagleson and Dharia, “Controlling FCCU emissions,” 11th Refin-
allowing it to be collected by the BELCO-F-Nozzle located at the top. ing Technology Meeting, HPCL, Hyderabad, 2000.
To ensure droplet-free stack, the flue gas enters a droplet separa-
tor. This is an open design that contains fixed-spin vanes that induce a Licensor: Belco Technologies Corp.
cyclonic flow of the gas. As the gases spiral down the droplet separa-
tor, the centrifugal forces drive any free droplets to the wall, separat-
ing them from the gas stream.
White oil and wax hydrotreating
��
Application: Process to produce white oils and waxes.
��
Feeds: Nonrefined as well as solvent- or hydrogen-refined naphthenic or ��
��
paraffinic vacuum distillates or deoiled waxes. ��
��
Products: Technical- and medical-grade white oils and waxes for plasti-
cizer, textile, cosmetic, pharmaceutical and food industries. Products are
in accordance with the US Food and Drug Administration (FDA) regula-
��
tions and the German Pharmacopoeia (DAB 8 and DAB 9) specifications. ��
��
Description: This catalytic hydrotreating process uses two reactors. Hydro- ��
��
gen and feed are heated upstream of the first reaction zone (containing a
��
special presulfided NiMo/alumina catalyst) and are separated downstream
of the reactors into the main product and byproducts (hydrogen sulfide ��
��
and light hydrocarbons). A stripping column permits adjusting product ��
specifications for technical-grade white oil or feed to the second hydro-
genation stage.
When hydrotreating waxes, however, medical quality is obtained
in the one-stage process. In the second reactor, the feed is passed
over a highly active hydrogenation catalyst to achieve a very low level
of aromatics, especially of polynuclear compounds. This scheme per-
Installation: Four installations use the Uhde (Edeleanu) proprietary tech-
nology, one of which has the largest capacity worldwide.
mits each stage to operate independently and to produce technical- or
medical-grade white oils separately. Yields after the first stage range Licensor: Uhde GmbH.
from 85% to 99% depending on feedstock. Yields from the second
hydrogenation step are nearly 100%. When treating waxes, the yield
is approximately 98%.

Utility requirements (typical, Middle East Crude), units per m3 of feed:

1st stage for 2nd stage for Food-grade
techn. white oil med. white oil wax
Electricity, kWh 197 130 70
Steam, LP, kg 665 495 140
Water, cooling, m3 48 20 7
Hydrogen, kg 10.0 2.6 1.6
Annual Meeting
March 19-21, 2006
Grand America Hotel
Salt Lake City, UT

AM-06-27 Integrated Hydrogen Solutions: Combining

Hydrogen Recovery and Optimized Steam
Reforming

Presented By:

James D. Fleshman
Senior Principal Engineer
Foster Wheeler USA
Corporation
Houston, TX

Mario Campi
Process Manager
Foster Wheeler
Milan, Italy

National Petrochemical & Refiners Association 1899 L Street, NW 202.457.0480 voice

Suite 1000 202.429.7726 fax
Washington, DC www.npra.org
20036.3896
Introduction

The design of a hydrogen supply system involves much more than specifying a plant to
produce a particular amount of hydrogen. In order to remain competitive in today’s
economy, it is necessary to optimize how hydrogen is produced, recovered and
compressed, particularly for a complex refinery.

Tighter regulations on sulfur content of transportation fuels continue to require additional

hydrogen capacity for hydroprocessing, while at the same time the regulations on
aromatic content of gasoline limit catalytic reformer throughput and byproduct hydrogen
production. The trend towards the processing of heavier crudes and the more
widespread use of residue conversion technologies further aggravates the hydrogen
shortfall. As a consequence the demand for hydrogen continues to increase.

This article addresses the methods for meeting these challenges and focuses on an
innovative, cost-effective solution developed by Foster Wheeler for the NIS Pancevo
Refinery in Serbia.

AM-06-27
Page 1
Hydrogen Management

The most economic solution to a hydrogen supply problem is to make best use of
existing resources. Adding a new unit to supply the entire shortfall without
optimization is rarely the low cost solution.

Developing an optimum supply scheme requires an evaluation of the hydrogen

demand, the configuration of the existing system, and available options for
increasing capacity. Foster Wheeler applies a proven approach to hydrogen
management, which addresses:
• The current refinery balance
• Hydrogen recovery from various refinery offgases
• Upgrading or revamping options for existing hydrogen production units.
• Additional new capacity
• Purchasing hydrogen over-the-fence

The final evaluation also factors in capital and operating cost, as well as
operability and reliability of supply.

AM-06-27
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Hydrogen Balance

When analyzing hydrogen resources in refinery complexes, Foster Wheeler uses

a hydrogen pinch technique to identify the best opportunities for offgas re-use
and purification and to quantify the minimum size of any new hydrogen
production facilities that may be required. The pinch technique aims to develop
the optimal hydrogen recovery scheme based on capital cost, operating cost, and
constructability. Key elements of the pinch analysis are:

• Confirmation of existing network headers

• Compression requirements
• Potential for re-use of offgases containing hydrogen
• Potential application for hydrogen purification
• Cost analysis

Revamp Opportunities

There are a number of options for upgrading existing hydrogen production units.
Depending on the age of the unit, new technologies are available in virtually all
areas of the plant. These include improved alloys for reformer catalyst tubes,
improved catalysts, and more efficient CO2 removal solutions or Pressure Swing
Adsorption (PSA) adsorbent.

Increased throughput can also generally be achieved by additions to the plant

such as a pre-reformer or post-reformer, or a LT shift converter.

An increase in capacity also requires changes to the remainder of the system,

including gas cooling and hydraulics, and at some point additional capacity
becomes prohibitively expensive. Other solutions for consideration involve
sharing of hydrogen supplies with neighboring facilities or an over-the-fence
supply from an industrial gas supplier. Once these options have been

AM-06-27
Page 3
exhausted, and a new hydrogen production unit identified, the new facilities need
to be integrated into the existing system.

Integrated Offgas Utilization

Because of economics, there has been an increased focus on the best way to
integrate available hydrogen-rich refinery offgas streams into hydrogen
generation capacity. Aside from the actual hydrogen balance, it is necessary to
consider purification, compression, and the impact on the refinery fuel balance.

Hydrogen expansion projects generally include offgas purification by PSA in

order to meet purity specifications. A new hydrogen plant will typically include a
Steam Methane Reformer (SMR), also with purification by PSA. In that case
there are typically three PSA configurations to be considered when developing an
integrated hydrogen supply system:

• Combine refinery offgas (ROG) with SMR gas upstream of a common PSA.
• Combine the ROG with feedstock upstream of the SMR.
• Send the ROG to a dedicated hydrocarbon based PSA unit where tail gas is
recompressed for either fuel export or utilized as reformer feed.

The first configuration, use of a common PSA unit, often appears the simplest
and can at least theoretically offer a cost-effective solution. There are, however,
a number of adsorbent and fuel balance problems with this configuration, which
limit its application.

The impurities present in SMR gas – CO2, CO and N2 – require a different set of
PSA adsorbents than are required by the hydrocarbons present in ROG. A PSA
unit with mixed feedstock requires adsorbents to remove both sets of impurities.
In that case there is a risk that the heavier hydrocarbons from the refinery gas
will reach the adsorbents for SMR gas and be permanently adsorbed, damaging

AM-06-27
Page 4
the bed and reducing its capacity. This makes the combined system very
sensitive to the flow and to the composition of the ROG, and puts the system at
risk of permanent damage if the flow or composition of the ROG varies.

This results in the following constraints on the combined PSA system, which
limits it usefulness:
- The ratio of ROG to SMR gas needs to be kept constant for optimum bed
usage without risk of bed contamination.
- If the SMR is shutdown, the PSA operation on ROG alone may result in
adsorbent contamination by heavy hydrocarbons moving too high in the bed.
- In many cases the offgas stream often contains less H2 and significantly
higher concentrations of heavy hydrocarbons than originally predicted. This
is difficult to foresee and can lead to irreversible adsorbent de-activation if it
goes undetected and adequate measures are not taken to adjust the PSA
cycle time.
- Because of the heavier hydrocarbons, the combined PSA tail gas has a much
higher calorific value, which is often in excess of the steam reformer fuel
requirements. A tail gas compressor is then required to route the excess gas
into the fuel gas network. Moreover, the excess tail gas contains CO2 from
the SMR, which can upset the refinery fuel system.

If the amount of ROG is small, it may still be feasible to combine SMR and refinery gas
in a single PSA unit. However, as the quantity of hydrogen contained in the ROG
increases, it becomes more practical to either use the ROG as SMR feedstock, which
eliminates the tail gas problem, or to install a separate PSA to purify the ROG streams.

Where separate PSA units are used, tail gas from the SMR PSA is returned to the
SMR as a dedicated fuel stream, while tail gas from the ROG PSA unit can be
compressed and used in the refinery fuel system.

AM-06-27
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Case History – NIS Pancevo

Foster Wheeler recently had the opportunity to review various integration options at
the NIS Pancevo refinery in Serbia.

NIS-Petrol has two refineries in Serbia: one in Pancevo and the other in Novi Sad,
with a total installed capacity of 5 million metric tons/annum (MMTPA). The
Pancevo refinery is undergoing a development program aimed at:
• Production of gasoline and diesel in accordance with European standards (EU
2010+)
• Meeting domestic and European environmental protection standards
• Increasing conversion
• Generating export products
• Energy optimization
• Increasing and securing refinery profitability

Key refinery units in the development include a Continuous Catalytic Reformer

(CCR), Mild Hydrocracker (MHC), Naphtha and Diesel Hydrotreater (DHT) plus
support units including a new hydrogen plant.

The hydrogen demand for the new units was 64,000 Nm3/h of high purity hydrogen.
Two hydrogen-rich offgas streams were available, one from the CCR unit and a
purge stream from the new MHC/DHT unit. Table 1 shows the composition of the
CCR and MHC/DHT offgas streams. The hydrogen recovered from these two
streams was to be supplemented with ”on purpose” hydrogen produced in a new
SMR. Natural gas was available as primary feed to the SMR with LPG available as
supplemental feed.

AM-06-27
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Table 1
Hydrogen-Rich ROG Stream Properties

CCR Offgas MHC/DHT Offgas

Component
Volume % Volume %
H2 94.93 73.95
Methane 1.52 13.17
Ethane 1.41 5.36
Propane 0.99 3.73
Butanes 0.56 2.26
Naphtha 0.59 1.11
H2S 0.00 0.01
H2O 0.00 0.41
Total 100.00 100.00

Flow, kg-mols/Hr 1,057.8 340.0

Pressure, bar-g 18 20
Temperature, °C 38 45

AM-06-27
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The hydrogen available in the two offgas streams was significant - approximately
45% of the total hydrogen demand for the new facilities. Based on this, the following
possible configurations were evaluated for integration of the ROG purification with
the SMR:

• Case A – ROG tail gas to fuel (the base case)

• Case B – ROG tail gas to SMR feed
• Case C – ROG directly to SMR feed

Case A – ROG Tail Gas to Fuel

In this case the offgas from the CCR and MHC / DHT units is purified in a dedicated
ROG PSA unit, and the tail gas is compressed to the refinery fuel gas header
pressure of 6 bar-g. Part of this tail gas is used as supplemental fuel in the SMR,
with the remainder exported to the refinery.

Natural gas from battery limits is compressed and used as feed in the SMR. A
separate PSA unit is used to purify the SMR gas, and tail gas from the SMR PSA at
0.3 bar-g is used directly as fuel for the reformer, supplemented by tail gas from the
ROG PSA unit.

In each case a small amount of hydrogen is recycled from the PSA outlet to the
SMR feed purification section for the hydrogenation and removal of organic sulfur.
The flows shown here neglect this recycle.

AM-06-27
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CCR MHC / DHT
22,493 4650
16

22,800 15
ROG 64,166

Nm3/h of 100 % H2
PSA
0.3
Pressure, barg

6 41,366
4,343

17 35 45,963
NATURAL
SMR SMR
GAS
PSA
4,597

FUEL
FG to
Header
2,271 2,072

Figure I - CASE A - ROG Tail Gas to Fuel

AM-06-27
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Case B – ROG Tail Gas to SMR Feed
As in the first case, the refinery offgas is purified in a dedicated PSA unit. However
in this case the tail gas is compressed and sent to the SMR to be used as feed.
Since hydrogen in the tail gas is recovered as SMR feed instead of being burned as
fuel, it is not as critical to maximize the hydrogen recovery in the ROG PSA unit. A
higher tail gas pressure can then be used to reduce the requirement for tail gas
compression: 1.0 bar-g vs. the 0.3 bar-g used in Case A. This increased pressure is
particularly valuable in this case, since the tail gas must be compressed to the
natural gas pressure of 17 bar-g. Since the PSA unit performance is tied inversely
to the tail gas pressure, the increased pressure reduces the hydrogen recovery in
the PSA unit from 84 to 81.5%. Natural gas from battery limits is then mixed with the
tail gas from the offgas PSA, compressed, and sent to the SMR as feedstock. As in
Case A, the SMR product is then purified in a dedicated PSA, with the tail gas used
as SMR fuel. In this case the supplemental fuel for the SMR is natural gas.

AM-06-27
Page 10
CCR MHC / DHT
22,493 4,650
16
64,166
22,122 15
ROG
Nm3/h of 100 % H2
PSA
1.0
Pressure, barg

42,044
5,021

17 35 46,716
NATURAL SMR
GAS SMR
PSA
0.3 4,672
FUEL
NATURAL
GAS

Figure 2 - CASE B - ROG Tail Gas to SMR Feed

AM-06-27
Page 11
Case C – ROG Directly to SMR Feed

In this case, the ROG PSA unit is eliminated and offgas from the CCR and
MHC/DHT unit is sent directly to the SMR feed compressor together with the
natural gas. The tail gas from the SMR PSA at 0.3 bar-g is used directly as fuel for
the reformer, with natural gas as the supplemental fuel.

CCR MHC / DHT

22,493 4,650

Nm3/h of 100 % H2

Pressure, barg 15
64,166

17 35 71,295
NATURAL SMR
GAS SMR PSA
7,129

FUEL
NATURAL
GAS

Figure 3 - CASE C - ROG Directly to SMR Feed

AM-06-27
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Evaluation of the Different Cases

The assessment of the three cases, A, B and C, took into account expected
availability as well as capital and operating costs.

The direct utilization of the ROG as SMR feedstock in Case C eliminates one set of
compressors and a PSA unit. This provides lower complexity and fewer processing
steps, resulting in improved operability. Complexity of the other two cases is similar.

From an availability viewpoint, Cases A and B are preferable since high purity
hydrogen can be generated when either the ROG PSA or the SMR system are
unavailable.

The non-availability of ROG has approximately the same impact in all three cases,
particularly since adequate margin was built in to the design of the SMR to partially
compensate for the loss of ROG.

A loss of the ROG PSA tail gas compressor impacts only Cases A and B.

The various operational aspects are summarized in Table 2.

AM-06-27
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TABLE 2
Operating Issues

Issue A B C
Complexity More More Less
H2 Availability High High Normal
No availability of ROG 34% loss of H2 35% loss of H2 38% loss of H2
Loss of ROG PSA tail
Loss of fuel 10% loss of of H2 No impact
gas compressor
Loss of feed gas
66% loss of H2 65% loss of of H2 Total loss of H2
compressor or SMR

Table 3 shows the differences in the capital cost of the three cases using Case A as
the base case. The costs are referred to the first quarter 2005, Western European
basis.

Table 3
Capital Cost Differences

$MM

Case A B C
Feed Compressors --- +0.1 +0.1
Steam Reformer --- --- +8
SMR PSA --- --- +1
ROG --- -0.6 -2
Tail Gas Compressors --- +0.2 -1
Total Cost Difference --- -0.3 +6

AM-06-27
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Operating Costs

In order to provide a consistent basis, the following operating parameters were

retained for each case:
• Steam to carbon ratio (3.0)
• Reformer inlet / outlet temperature (650/880°C)

The main process parameters - the resulting feed and fuel consumption, export
steam flows and related operating costs - are reported in the following tables.

TABLE 4
MAIN PROCESS PARAMETERS

A B C
H2 Recovered from ROG Nm3/hr 22,800 22,122 -
H2 Produced from SMR Nm3/hr 41,366 42,044 64,166
Total Hydrogen Production Nm3/hr 64,166 64,166 64,166

Reformer Absorbed Duty Gcal/hr 33.91 32.23 40.99

SMR PSA Efficiency % 90 90 90
ROG PSA Efficiency % 84 81.5 -

TABLE 5
FEEDSTOCK AND UTILITY CONSUMPTION

A B C
Natural Gas Feed Gcal/hr 131.2 69.0 90.8
Natural Gas Fuel Gcal/hr - 27.3 13.4
ROG Feed, Gcal/Hr Gcal/hr 115.6 115.6 115.6
ROG Tail Gas Export Gcal/hr 27.9 - -
HP Steam Export MT/hr 34.7 34.6 40.0

AM-06-27
Page 15
TABLE 6
FEEDSTOCK AND UTILITY COSTS - $MM/YR

A B C
Total Feed Plus Fuel
(Natural Gas plus ROG, 28.58 24.90 26.33
less ROG Tail Gas Export)
HP Steam Export (6.24) (6.23) (7.20)
Other
(Electricity, Demin water, 3.32 3.51 3.88
cooling water, LP steam)

Total 25.66 22.18 23.01

In Summary
• Case A and B show lower investment cost
• Case B and C show lower operating cost
• Case B shows the lowest combination of investment and operating cost

Case B was therefore selected.

Hydrogen Plant Design

In addition to developing the system configuration, the project also included design
of the hydrogen plant to meet the objectives of NIS Pancevo, including:
• Flexibility to use natural gas or LPG feedstock
• Efficient and reliable hydrogen supply
• Variable cost as low as possible without incurring excessively high
capital expenditure
• Ability to produce export steam cost effectively

NIS Pancevo’s main objectives were met by the installation of a Foster Wheeler
Terrace Wall® Steam Reformer with microalloy catalyst tubes and state-of-the-art
catalyst.

AM-06-27
Page 16
Because of the requirement for feedstock flexibility, a pre-reformer was integrated
with the Foster Wheeler Terrace Wall® primary reformer. This allowed extended
operation on heavier feedstocks, as well as improved energy efficiency. Since
heavier hydrocarbons are not present in the effluent from the pre-reformer, this
stream can be preheated to a higher temperature before it enters the reformer
radiant section. This reduces the load on the reformer radiant section and reduces
fuel requirements.

The other main features of the Terrace Wall ® furnace which insured that NIS-
Pancevo’s targets would be met are as follows:

1. The Terrace Wall ® design uses a single row of catalyst tubes located inside each
radiant cell. Specially selected burners are arranged in terraces at two levels along
a firing wall to create a nearly uniform heat flux to the double-fired tubes. Slightly
sloping walls assist in providing uniform heat distribution since they slope towards
the tubes as the gas cools. Since the burners are at two levels, flux adjustments can
be made along the tube length as the catalyst ages.

2. The entire Terrace Wall reformer, including radiant and convection sections, can be
brought to the site in modules. This reduces field labor as well as construction cost
and duration. These modules can be shipped over the road to virtually any location.

3. Since the selected reformer at Pancevo used natural draft, no fans were required
and reliability was improved. The Terrace Wall furnace also allows the ability to
install combustion air preheat, and switch to natural draft operation in case of fan
failure.

4. Burners are arranged in two firing levels, which allow optimization of the heater
efficiency and the hydrogen production in accordance with the type of feed, catalyst

AM-06-27
Page 17
life and unit load. There is virtually no limitation in turndown and the steam reformer
can maintain a high efficiency at reduced load.

5. The isolation and replacement of radiant tubes are simplified, since the inlet pigtails
are fully accessible while the outlet pigtails are easily accessible through removable
panels on the outlet manifold box. Each single tube can be removed from the top
without any additional work except the removal of the roof rain cover. In case of
tube failure the tube can be crimped and isolated with the heater in operation.

6. Loading or unloading of the catalyst requires only the removal of the top flange.
Loading of a complete charge of catalyst can be completed in two to three days.

In order to supply the maximum amount of steam to the refinery, combustion air preheat
was not selected and heat recovery was by maximizing steam generation. The Terrace
Wall® design, with the convection section placed on the top of the radiant section,
allowed natural draft operation. This eliminated the induced and forced draft fans,
which further improved the expected availability and reduced operating costs.

Foster Wheeler’s combination of hydrogen management and process expertise,

coupled with the expertise of Foster Wheeler Fired Heater Division, provided NIS
Pancevo with a cost-effective and highly reliable solution for meeting the hydrogen
demands associated with upgrading the refinery to a state-of-the-art facility.

AM-06-27
Page 18
This paper has been reproduced for the author or authors as a courtesy by the National
Petrochemical & Refiners Association. Publication of this paper does not signify that the
contents necessarily reflect the opinions of the NPRA, its officers, directors, members, or staff.
Requests for authorization to quote or use the contents should be addressed directly to the
author(s)
Reprinted from:
June 2006 issue, pgs 57 –62
Used with permission.
SPECIALREPORT
www.HydrocarbonProcessing.com
Article copyright © 2006 by Gulf Publishing Company. All rights reserved.

Upgrade refinery residuals

into value-added products
A European refiner effectively uses delayed coking
to exit residual fuel-oil market
M. KOVAC and G. MOVIK, MOL Plc Danube Refinery, Szazhalombatta,
Hungary; and J. D. ELLIOTT, Foster Wheeler USA Corp., Houston, Texas, US

T
he MOL Plc Danube refinery in Szazhalombatta,
Hungary, faced several operations challenges: 1)
declining residual fuel-oil markets and 2) a need to
increase clean transportation distillates output. The refinery
considered a residue upgrading project that would increase
conversion of crude oil to distillates, eliminate residual
fuel oil yields and improve refinery margins. After study-
ing alternative paths, MOL selected delayed coking as the
primary conversion process.
In 1997, the MOL board resolved to implement a resi-
due upgrading project based on delayed coking. This proj-
ect also included new facilities for a delayed coker, hydro-
gen plant and sulfur recovery unit, as well as, revisions to
the distillate hydrotreater and conversion of a vacuum
gasoil (VGO) hydrotreater to a licensed mild hydrocracker. FIG. 1 The MOL Plc Danube refinery, Szazhalombatta, Hungary, with the
new delayed coker complex.

specific situation. Delayed coking is the preferred choice for

Driving forces for change in refining many residual conversion projects. This processing method
Refiners are under more pressure to improve profitability offers several benefits:
while maintaining or increasing market share. Driving forces • Complete conversion of residue feedstocks and elimina-
within the refining industry include: tion of residual fuel-oil production. Demetallization is
• Increasing demand for clean, distillate nearly 100%.
transportation fuels • Ability to process very heavy residue streams with high
• Higher crude oil costs, especially for light, sweet crudes metals, carbon residue and asphaltenes
• Declining heavy residual fuel-oil markets. • Operational flexibility to handle a broad range of feed
Residual-oil conversion projects can be effectively used to qualities and adjustable product specifications
meet these needs. Primary conversion processes commonly • Ability to produce a broad range of liquid distillates that
evaluated for such purposes are: can easily be incorporated into the refinery processing
scheme to produce clean distillate transportation fuels
Noncatalytic
• Moderate capital investment
• Visbreaking
• The process has been commercialized for a long time and
• Solvent deasphalting
is well supported by specialty equipment vendors, chemical
• Delayed coking
suppliers and third-party consulting firms
• Flexicoking or fluid coking
• Byproduct fuel-grade coke is easily salable at positive net-
• Partial oxidation back prices
Catalytic • Modern delayed coker designs are proven to be energy
• Ebullated-bed hydrocracking efficient, environmentally friendly and intrinsically safe
• Slurry-phase hydrocracking • The semi-batch nature of the delayed coker facilitates
• Residual fluid catalytic cracking. easier operations and provides schedule flexibility due to
All of these processes have good features and specific appli- equipment failures
cations. Thus, the refiner must evaluate these options and con- • Operating and maintenance costs are reasonable, and peri-
version routes, and select the best method for the refinery’s ods between turnarounds can be as long as five years.

HYDROCARBON PROCESSING JUNE 2006

SPECIALREPORT PROCESS AND PLANT OPTIMIZATION

��
��
��
��

FIG. 3 Product yields of the delayed coker-design vs. actual.

��

FIG. 2 New delayed coker, side view.

The project goals were:

• Provide an exit from the uncertain heavy fuel-oil market
• Strengthen competitiveness and increase crude conversion
level ��
• Improve environmental conditions. ��
Proposals from qualified delayed coker licensors were solic- ��
ited and evaluated. An experienced delayed coking technology
licensor was selected. The technology licensor would provide FIG. 4 Variance in sulfur levels of coker products-design vs.
the delayed-coker technology, supply the coker heater, and sup- actual.
port training and commissioning efforts for unit startup. Detail
engineering and construction was to be provided by a local
contractor with support from the technology licensor for coker-
specific issues (Fig. 1).
��

Design basis coker. The coker was specified to process 3,250

m3/d (20,442 bpsd) of blended vacuum residue. In addition to
vacuum residue, the coker was designed to process other extrane-
ous streams:
• 7,000 metric ton/yr (mtpy) of refinery sludge on an
intermittent basis
• 26,400 mtpy of propane/propylene from the FCC unit.

Products obtained from the delayed coker included: ��
• Treated coker fuel gas (FG) ��
• Propylene (99.5% purity) ��
• Propane
• Treated C4 LPG FIG. 5 Specific gravities of coker products-design vs. actual.
• Light coker naphtha (LCN)
• Heavy coker naphtha (HCN)
HYDROCARBON PROCESSING JUNE 2006
SPECIALREPORT PROCESS AND PLANT OPTIMIZATION

TABLE 1. Design features of the delayed coker TABLE 3. Coke quality—design vs. actual

• Advanced coke drum design for optimum life. Specifications included drum Design feed prediction Actual feed operation
plate chemistry and thickness and fabrication instructions. VCM, wt% 8 –11 10.2
• Side-wall double-fired coker heater fitted for online spalling. This design Ni+V ppmw 1,026 1,195
provides long heater run length by minimizing reaction time in the heater.
Heating value, kJ/kg 35,674 35,800
• Safety interlock systems for coke drum isolation, coke cutting system and
HGI 50– 80 64
heater operation. An enclosed operator shelter was specified for the decoking
operation on the top deck of the coke drum structure. Moisture, wt% 12– 15 9.3
• The heater was supplied with air preheat and low NOx features
• Enclosed blowdown system for recovering unconverted oils and vapors TABLE 4. Refinery products before and after coker
without impacting the environment. Result: No scheduled flaring due to drum project
switching and turnover
• Decoking water systems designed for 100% recovery and recycling of water Before After
coker, % coker, %
• Sludge injection system to process refinery sludges and convert them to
product Gasoline 22.4 24.0

• Gas plant with 90%+ recovery of C3 streams; amine scrubbing of coker Naphtha 6.8 8.4
product gas prior to refinery export and fuel use in the coker heater; amine LPG 4.5 5.3
treating of the C3/C4 LPG stream and a naphtha splitter Middle distillates 37.7 42.7
• Licensed sections from others include a C3/C4 LPG caustic treater and Heavy fuel oil 14.6 1.2
propane-propylene splitter
Coke 3.4
Other 14.0 15.0
TABLE 2. Feed quality variance 100.0 100.0

Specific gravity –.0085

Carbon residue, wt% +2.3
to provide assistance with precommissioning, unit checks, vendor
coordination, onsite training and operating instruction support.
• Light coker gasoil (LCGO) In mid-summer, he was joined by three additional operations
• Heavy coker gasoil (HCGO) advisors and an instrument engineer to complete the commis-
• Fuel-grade coke. sioning effort.
Unit startup and test run. Following strong client-licensor
Features of delayed coker design. Processing objectives teamwork in preparing for the startup, the unit was started up on
dictated that the delayed coker operating conditions be configured Oct. 6, 2001, without incident at 80% capacity. The unit capacity
for low coking pressure at 1.03 BarG in the coke drum and ultra- was increased to 100% on October 10 with all products “on spec.”
low recycle ratio at 5% liquid volume to produce maximum clean About three weeks later, the coke-drum switch valve experienced
distillates and minimize fuel-grade coke production (Fig. 2). With an unforeseen and unusual metallurgical failure. While the valve
this style of operation, coke production was limited to 230,000 was repaired by the vendor on an expedited basis, the unit was
mtpy. The green coke was slated for sale to the local power and idled for two weeks and restarted on November 14. A capacity
cement industries. Table 1 summarizes the design features for the test at 110% was completed on November 22. The performance
delayed coker. guarantee test was done on December 6, 7 and 8. All guarantees
During detail design, MOL elected to add several features: were met.
• Advanced hydraulically operated coke drum unheading
systems Delayed coker performance. During evaluation, it was
• Advanced coke cutting system with hydraulically operated determined that the coker feedstock was higher in carbon residue
motors for the drilling system than the design feed. Table 2 summarizes the variance of the feed
• Enclosed coke transport and storage using environmentally and change in quality.
friendly tubular conveyors and a coke storage “barn” A comparison of the design and evaluation yields is shown in
• Advanced process control and operator simulation through Fig. 3. The design yields and product properties were based on
the DCS. a proprietary model for delayed coking. Pilot-plant runs were
Commissioning and startup. The preparation for startup not required. The resulting yields and product properties had
involved a team effort between MOL and technology licensor in small variances from the general model after feed differences are
various activities. considered.
Operating guidelines. The licensed package was supplied Regarding the actual operating yields, note the higher coke
with operating guidelines from which MOL was able to develop yield as a function of the heavier-than-design feedstock. One
their operating instructions. comment on the yield pattern was a displacement of light coker
Training. During the summer of 2000, MOL was provided gasoil for heavy coker naphtha experiences with this feedstock.
with classroom training and additional third-party training at an This pattern was investigated and has subsequently been incorpo-
existing delayed coking facility. rated into the proprietary delayed coker yield models.
Precommissioning and commissioning. In May 2001, the tech- Fig. 4 shows a relative comparison of the expected sulfur dis-
nology licensor’s lead operations advisor was sent to Szazhalombatta tribution and actual values. The variances were acceptable to the
HYDROCARBON PROCESSING JUNE 2006
SPECIALREPORT PROCESS AND PLANT OPTIMIZATION

hydrotreating units. The actual sulfur distribution is considered during the design phase. This project is profitable and allowed
environmentally favorable. Fig. 5 illustrates the differing gravities MOL to economically reach its goals of exiting the uncertain
of the liquid products. Table 3 lists the design estimate for coke heavy fuel-oil market, improving competitiveness and improving
quality with the actual coke. environmental conditions. The success of the coker operation
Coke disposal is 30% to local steel and cement industries and is due to an effective teamwork between refiner and licensor to
70% export to purchasers in Europe. After accounting for vari- define the design, control the execution and start up the unit
ances in operation and feed quality, the design basis for yields and without incident. HP
product quality was confirmed by the operating data.

Impact of coker on operations and profitability. As Miroslav Kovac is a supervisor of the delayed coker unit
in MOL Plc’s Danube Refinery. He has 12 years of experience in
a result of the coker project, the MOL Szazhalombatta refinery petroleum refining in the fields of process engineering and opera-
was able to effectively exit the heavy fuel-oil market. Table 4 lists tion. Previously, he worked as a research fellow at the University of
the percentage shift in refinery products before and after the coker Veszprem. He holds an MS degree in chemical engineering from
project. the University of Veszprem.

The profitability of the delayed coker is high. In the second-

quarter 2002 statement of financials, MOL stated:
Gabor Movik is a process engineer of the residue processing
“The operating profit of the refining and marketing area in MOL Plc’s Danube Refinery. He has seven years of experi-
segment increased by 9% to HUF 21.1 billion compared ence in petroleum refining in the field of process engineering. He
to the second quarter last year, in spite of the lower crack holds an MS degree in chemical engineering from the University
spreads and reduced Brent-Ural differential. . . The start of Veszprem.
of the delayed coker increased further the efficiency of the
refining process and contributed HUF 3.1 billion to group
performance.” John D. Elliott is director, refining and coking, for Foster
Wheeler USA Corp., Houston, Texas. He joined Foster Wheeler in
Based on the currency exchange at the time, this is equivalent to 1967 and has over 30 years’ total experience in refining process
€140,000/day (US$129,000/day). engineering. He is recognized as a leading authority on heavy oil.
Mr. Elliott’s assignments have involved process design and operating
follow-up on a number of refining units, including over 50 major delayed coker proj-
Outlook. MOL has executed a very successful coker project at
ects. Mr. Elliott has presented many papers on heavy oils processing and, in particular,
the Szazhalombatta Refinery. The delayed coker is the center- delayed coking, at major industry seminars globally. He holds a BS degree in chemical
piece of the projects and has achieved the processing targets set engineering from Pennsylvania State University, and is a member of AIChE.

Article copyright © 2006 by Gulf Publishing Company. All rights reserved. Printed in U.S.A.
Not to be distributed in electronic or printed form, or posted on a website, without express written permission of copyright holder.

2020 Dairy Ashford

Houston, Texas 77077
Phone 281.597.3000
Fax 281.597.3028
Email coking@fwhou.fwc.com
www.fwc.com
DRIVERS FOR ADDITIONAL DELAYED COKING CAPACITY IN THE REFINING INDUSTRY

The refining industry has for some time been facing weak and volatile refining margins and asset
expansion is likely to remain limited. Refiners have historically had access to very limited capital
expenditure budgets with investment limited to ‘stay in business’ expenditure, generally driven by the
need to meet increasingly stringent product quality legislation. Certainly, where more strategic
investment has proceeded, companies have been under pressure to achieve shorter payback periods.

Bottom-of-the-barrel processing is increasingly likely to be considered, certainly in USA and Europe, for
a number of reasons:

• Growing demand for transportation fuels ) The combination of these leading to an

• Potential reduction in markets for residues ) unbalanced production slate at refineries
• Crude oil quality, availability and light/heavy price differentials
• Refineries’ competitive position
• Demand for anode grade coke
• A reluctance to meet demand changes by investment in additional grassroots crude processing

Growing transportation fuel demand

Global transportation fuels demand is significantly influenced by global events affecting economic
activity and confidence. Hence demand growth has historically been erratic. World consumption of
refined products grew by an annual rate of 1.4% during the 1990s. However, if the Former Soviet Union
and Eastern Europe are excluded, the growth rate is higher at 2.4% per annum.

In this period all light products (transportation fuels and petrochemical intermediates) exhibited
stronger growth, with global demand expected to average approximately 2% per annum. We are seeing
a shifting demand/supply balance, especially for middle distillates. In Europe, for example, the deficit
for road diesel is forecast to be around 45 million tonnes by 2010.

Figure 1: Road Diesel Deficit

Source: TotalFinaElf, WEFA Conference, June 2001
0

-5

-10
million tonnes per year

-15

-20

-25

-30

-35

-40

-45
2000
2005
2010

Page 1 of 4
Residual markets

For most refiners fuel oil is usually an undesirable byproduct. Unlike light products (such as gasoline
and middle distillates), fuel oil usually commands a price below the cost of crude thereby effectively
depressing overall refinery margins.

Residual fuel oil demand declined by 20% during the 1990s even though bunker fuel use increased over
the same period. At a recent bottom of the barrel conference (ref 1) Nexant/Chem Systems reported
that a global 50% fall in residual fuel markets would necessitate around 200 heavy residue destruction
projects with current crude slates. In Europe, legislation regarding the sulfur content of fuel oil may
reduce the market size by 50%, leaving bunker markets as the main outlet. If the pace of market
decline were to accelerate, many refiners might face ‘stay in business’ investment decisions,
converting residuals by means of capital intensive processes with unexciting rates of return.

Some forecasters, however, believe that residual fuel oil demand has now leveled out and could even
increase by around 0.5% per annum. This is being caused by a combination of US national gas
shortages, the early phasing out of nuclear power in certain countries and growth in demand for
bunkers. However, this growth rate is four times lower than the forecast global growth in light product
demand.

Crude supply & price differentials

Directionally, world crude supply is expected to become heavier and more sour.

Light crude (34 API or higher) represents around 50% of crude currently processed. Although production
of light sour crude is expected to increase by 9 million b/d by 2015, the production of light sweet
crude is expected to increase by only 1-2 million b/d over the same period.

The pace of decline in light sweet availability will be influenced by production levels in West and North
Africa, the North Sea and elsewhere. The availability of light condensates could also have a major
influence on light feedstock supplies. Refiners have to some degree been able to avoid investment in
upgrading by crude substitution, especially in market conditions in which reduced light/heavy price
differentials have prevailed. However, fuel specifications are becoming so stringent that this solution is
not sufficient to meet the new specifications and residue upgrading is required.

Heavy crude release to the market has been held back by adherence to OPEC quotas (which encourage
sales of higher price light crudes). Heavy sour crude makes up only a small percentage of current
production but is increasing.

Light/heavy crude differentials can be explained largely by two key factors - supply, estimated by the
relative proportions of these grades in the output slate (the supply side), and demand for heavy
products, particularly residual fuel oil (the demand side). These factors pull in different directions,
with spreads widening as the relative output of heavy crude rises and narrowing with stronger fuel oil
demand. There is no reason to suspect that the volatility in differentials will disappear and
directionally, one would expect differentials to widen.

Page 2 of 4
Much of the investment in coking in the USA in the past 10-15 years has been driven by the wish to
process lower value crudes from Latin America. In broad terms about 50% of the potential economic
benefit from a coker comes from using cheaper crude (the other 50% comes from the higher value of
the upgraded products).

The bigger the spread between heavy and light products, the more incentive refiners have to invest in
the necessary hardware to process cheaper feedstocks to gain competitive advantage. The heavy
crude oils produce a higher portion of heavy, highly contaminated materials, which historically have
been disposed of to fuel oil. However, these stocks can be upgraded to more valuable products using
mature technologies such as fluid catalytic cracking (FCC), hydrocracking and delayed coking, thereby
helping to correct finished product imbalances to improve refinery margins. Typically, a delayed coker
will convert 65% of its feed into transportation fuels, the other 35% being coke.

Coke market

Fuel grade coke is an internationally traded product and competes with coal in the market place. Its
higher heating value over coal (typically 20% higher) can encourage its use, with traditional power
stations able to blend around 20% into coal feed without the need for significant modifications. The
high sulfur content of fuel grade petroleum coke usually means that it can only be used on power
plants with flue gas desulfurisation. Coke traders consider that the future market for fuel grade coke
is strong although it is generally regarded as a low value, distress product.

Anode grade coke, used in aluminum smelting, is a robust market and fetches a far higher price than
fuel grade ($300+ per tonne). To a large extent the production of anode grade coke is feedstock driven
requiring the processing of a low sulphur, low metals crude. Overall refinery economics, in general,
dictate whether purchasing this higher quality, more expensive crude, to make anode coke is
justifiable.

The demand for needle coke, used in the steel industry, has been fairly flat over the last 20 years. It is
unlikely any new needle cokers will be built for economic reasons.

The delayed coking option

Delayed coking is a mature process which remains for many the preferred residue upgrading option
because of its ability to handle the heaviest contaminated crudes. Globally, approximately one third of
installed residue upgrading plant is by delayed coking.

Today, around 50% of the worldwide delayed coking capacity is concentrated in the US with more than
2,000,000 B/D of installed capacity. In the last 15 years, delayed coking capacity has grown by 56% in
the USA, followed by hydrocracking (37%) and FCC (14%).

A recent paper by Foster Wheeler (Ref 2) compared the economics of a number of residue upgrading
schemes:
• Delayed coking/fluid catalytic cracking
• Integrated Gasification Combined Cycle (IGCC)
• Atmospheric residue hydrotreatment and residue catalytic cracking (RFCC)

In summary, if transportation fuels command a higher value to the refiner than power, the refiner is
most likely to opt for a delayed coking or RFCC-based scheme.

Schemes based on FCC/delayed coking, hydrocracking/delayed coking appeared to provide the lowest
capital and operating costs. All three schemes offered a positive internal rate of return.

Page 3 of 4
Coking offers refiners much more flexibility in feedstock and crude selection, allowing the refinery
scope to take advantage of spot markets. Coking offers the refinery the opportunity to move towards
zero fuel oil production whilst meeting future product demand growth from low-value residuals rather
than additional crude processing.

Ref 1) EPC 2nd Bottom of the Barrel Technology Conference, Istanbul, 8-9 October 2002.
Keynote address by Nexant/Chem Systems.

Ref 2) Advances in residue upgrading technologies offer refiners cost-effective options for
zero fuel oil production. Graham Phillips, Fang Liu, Foster Wheeler 2002 European
Refining Technology Conference, Paris.

BY:

Graham Phillips
Technology Manager Refining, Foster Wheeler

Michael Stewart
Senior Planning Consultant, Foster Wheeler

Randy Wolf
Regional Vice President, Foster Wheeler

This article has been published in Petroleum Economist, September 2003.

Page 4 of 4
WHEN SOLVENT DEASPHALTING IS THE MOST
APPROPRIATE TECHNOLOGY FOR UPGRADING RESIDUE
Edward J. Houde - UOP LLC: edward.houde@uop.com
Michael J. McGrath - Foster Wheeler USA: michael_mcgrath@fwhou.fwc.com

Introduction present and future users of solvent deasphalting

Solvent deasphalting has a key role in today’s technologies.
refinery, as the technology can be used in a
variety of uses for residue upgrading. It is less The merging of Foster Wheeler’s and UOP’s
expensive to build and operate than other residue experience in designing solvent deasphalters
conversion processes and is especially useful in represents a total capacity of more than 500,000
recovering large quantities of high quality oils BPSD and more than 50 commercial units with
which can be further upgraded via traditional FCC capacities ranging from 500 to 45,000 BPSD.
and Hydocracking units. These commercial applications have included all
applications of solvent deasphalting such as:
This paper outlines the benefits of the combined ∗∗ the production of lube oil feedstocks
UOP/FW Solvent Deasphalting (UOP/FW SDA)
offerings. The technology applications, process
∗∗ the recovery of incremental feedstock for
variables, key design issues, and technical
downstream FCC and hydrocracking units
support services are also discussed.

The fit for a particular location is particularly ∗∗ the production of road bitumen, and have
dependant on finding a use for the SDA pitch and included both two-product (deasphalted oil
various options are discussed. (DAO) and pitch) and three-product
(deasphalted oil, resin and pitch) process
configurations.
Background
While solvent deasphalting (SDA) has been used
for more than fifty years to upgrade non-volatile This merging of FW and UOP solvent
residues, the technology continues to evolve over deasphalting technologies provides our clients
time. It is a robust economical process that uses with the widest range of experiences, process
an aliphatic solvent to separate the typically more features, engineering know-how, technical
valuable oils and resins from the more aromatic support, and most importantly, a more efficient
and asphaltenic components of its vacuum residue and lower cost design.
feedstock. The earliest commercial applications of
solvent deasphalting used propane as the solvent Technology Overview
to extract high-quality lubricating oil bright stock
from vacuum residue. These applications were Solvent deasphalting, whether for the production
called propane deasphalting (or propane of lubricating oil or cracking stocks, uses a light
deresining when used to separate high molecular hydrocarbon solvent specifically tailored to ensure
weight resins from Pennsylvania-grade vacuum the most economical deasphalting design. For
residue). Solvent deasphalting process have example, propane solvent may be specified for a
gradually extended to include the preparation of low deasphalted oil yield operation such as lub
feedstocks for catalytic cracking, hydrocracking, production, while a solvent containing
and hydrodesulfurization units, as well as the hydrocarbons as heavy as hexane may be used to
production of specialty asphalts. obtain a high deasphalted oil yield for the
production of additional conversion unit feedstock.
In 1996, UOP and Foster Wheeler USA Plant designs have been developed using heavy
Corporation (FW) entered into a collaboration solvents at elevated temperatures in order to
agreement to merge their solvent deasphalting maximize the yield of usable deasphalted oil and
technologies. Both companies had extensive minimize the yield of pitch having a softening point
backgrounds in solvent deasphalting, and both of 350°F or higher.
companies had recently entered into collaboration
covering other residue upgrading technologies UOP/FW SDA technology is unique by different
such as visbreaking and delayed coking. A from other solvent deasphalting technologies in
collaboration covering solvent deasphalting was that it is not just one technology, but rather a
considered a logical fit that would benefit the combination of technology features and options

IDTC Conference
London, England Page 1 of 11
February 2006
WHEN SOLVENT DEASPHALTING IS THE MOST
APPROPRIATE TECHNOLOGY FOR UPGRADING RESIDUE
Edward J. Houde - UOP LLC
Michael J. McGrath - Foster Wheeler USA Corporation

which have been developed by both companies. by incorporating both stripping and
These features and options allow UOP/FW to rectification of the oil feed. Superior
offer the optimum process design for any solvent quality DAO is obtained from the RDC
deasphalting application. even at DAO recovery rates exceeding 85
percent, with an even greater difference in
For example, UOP has predominately focused its quality being seen at lower DAO yields.
technology on downstream conversion unit
applications. As a result, UOP’s solvent ∗∗ Supercritical solvent recovery. Supercritical
deasphalting experience has principally focused recovery of the solvent allows more efficient
on the use of butane and heavier type solvents utilization of the system’s thermodynamic
that can obtain higher DAO recoveries. There is a characteristics while also reducing the unit’s
distinct advantage to the use of supercritical operating costs.
separation for the recovery of the solvent and
DAO when using these types of solvents. ** Multiple product recovery designs that take
Consequently, UOP developed supercritical advantage of the changes in liquid-liquid
solvent-DAO separation technology. The other equilibrium that result from changes in
area that UOP focused its development efforts operating conditions between those utilized
involved minimizing the solvent to oil ratio while during extraction and those used for solvent
still producing a reasonably high quality DAO. recovery.

FW’s SDA technology development emphasis was ∗∗ Lower solvent requirements used to achieve
initially more focused on lower lift, very high quality processing objectives. Although increasing
applications, such as the production of lube oil the amount of solvent used in the extraction
feedstocks for hydrocracking and further solvent improves the extraction efficiency, it also has
refining. Consequently its technology originally a major impact on the unit’s operating costs.
focused around propane/butane deasphalting Consequently, the lowest solvent-to-oil ratio
using optimized extraction techniques for those necessary to achieve the desired product
specific applications. Additionally FW has made separation is typically specified.
available its detailed design and construction
experiences from a multitude of SDA projects. ∗∗ Optimal design of heat exchange systems.
UOP and FW’s combined design experience
Technology Advantages in optimizing SDA heat exchange systems
allows the SDA user to select a multitude of
The UOP/FW SDA technology has several distinct heat exchange options, depending on the
advantages that ensure the refiner obtains the project specific objectives and opportunities.
most efficient, economical, and flexible SDA
process. These include: Extraction Devices

∗∗ Novel extraction devices tailored to the The efficiency of the extraction process is the key
specific application: equipment design variable impacting both the
capital and operating costs of SDA. The
- UOP/FW/Sulzer’s Structured Packing and extractor’s role in SDA is to separate the
proprietary internals in both the multi- precipitate (pitch phase) from the continuous fluid
stage counter current extractor as well as stream (DAO/solvent).
the DAO and resin separators. This
technology provides state-of-the-art Both single-stage co-current extraction, where the
contacting and separation devices to bulk of the solvent is mixed with the feed prior to
maximize extraction efficiency as well as the extractor, as well as multi-stage counter-
optimal recovery of clean products. current extraction, where the bulk of the solvent
enters the bottom of the extractor separate from
- FW’s multi-stage rotating disk contactor the feed, have been used commercially.
(RDC). The RDC is specifically designed
to achieve high product yields and quality Structured packing or RDCs used in multi-stage

IDTC Conference
London, England Page 2 of 11
February 2006
WHEN SOLVENT DEASPHALTING IS THE MOST
APPROPRIATE TECHNOLOGY FOR UPGRADING RESIDUE
Edward J. Houde - UOP LLC
Michael J. McGrath - Foster Wheeler USA Corporation

counter-current extraction provide both the The impact of solvent-to-oil ratio on both capital
contacting area and time required for extraction as and operating costs is conservatively reflected in
well as segregation of the stages to reduce back Tables 2 and 3. These tables summarize an
mixing below the feed stage. Above the feed internal study undertaken by UOP and FW to
stage, structured packing or coils provide for the assess the relative costs of different solvent-to-oil
coalescing of entrained droplets of feed or pitch. ratios. The cost of fuel has significantly increased
since this study was performed and the savings at
Supercritical Solvent Recovery lower solvent ratios would be even more
pronounced. While the cost benefits at a lower
Although, often referred to as supercritical solvent ratio can be substantial, the solvent-to-oil
extraction, it is the solvent separation, not the ratio also has an impact on DAO quality (see
extraction that is carried out in the supercritical Figure 1). Consequently, the optimal solvent ratio
region of the solvent. The use of supercritical is determined based on the DAO’s downstream
solvent recovery results in a simpler process flow. processing requirements.
Gone is the need for multiple flash towers and
condensers associated with conventional multiple Figure 1: Effect of Solvent to Oil Ratio
effect evaporative type solvent recovery systems. 70
3/1 5/1

8/1
PPM Metals in DAO
60
Supercritical solvent recovery allows for more 50
efficient utilization of the system’s thermodynamic 40
characteristics. Presented below is a comparative 30
tabulation of the utility requirements of 20
conventional subcritical evaporative and 10
supercritical solvent recovery systems. Note that 0
the difference from subcritical to supercritical 20 25 30 35 40 45 50 55 60 65 70

indicates a shift from steam to fuel. When the DAO yield

latent heat content of the steam is considered, the

comparison indicates a significant reduction Optimal Design of Heat Exchange Systems
(20 percent) in heat input into the supercritical
system as compared to the subcritical system. From a heat exchange perspective, there are
This can translate into a total utility cost savings of several ways in which a SDA unit’s operating and
about 30 percent relative to sub-critical recovery. capital cost can be significantly changed. Careful
analysis of the application, its associated utility
costs, customer preferences, and process
Table 1 flexibility requirements need to be addressed to
Comparative utility requirements ensure the best design is applied for the specific
Sub- Super-
application.
critical critical Difference
For example, the heat exchange between the
Electricity, 1.2 2.0 0.8 DAO/solvent phase and the solvent will have an
kWh/bbl Feed optimal temperature of approach. Both UOP and
Steam, lb/bbl 125.0 12.2 -112.8
FW have world-class experts in pinch analysis that
Feed will ensure the SDA unit’s design optimizes the
tradeoff between capital and operating costs.
Fuel, M Btu/bbl 55.0 115.0 60.0 Other areas where heat exchange costs can be
Feed significantly altered are with heat exchange of the
pitch and DAO products.
Lower Solvent Requirements
This is discussed in further in detail in the Utility
Because UOP/FW SDA has the advantage of Requirement section of this paper.
highly efficient extraction technology, the solvent-
to-oil ratio can be minimized for the same
objectives of DAO yield and quality.

IDTC Conference
London, England Page 3 of 11
February 2006
WHEN SOLVENT DEASPHALTING IS THE MOST
APPROPRIATE TECHNOLOGY FOR UPGRADING RESIDUE
Edward J. Houde - UOP LLC
Michael J. McGrath - Foster Wheeler USA Corporation

Table 2
UOP/FW SDA Advantage
Capital Costs
Option: 1 2 Delta
Case: High Solvent to Low Solvent to _
Oil Ratio Oil Ratio
Solvent-to-oil Ratio (S/O) 8 5 3
Solvent C4 Mix C4 Mix -
Size (BPD) 28,000 28,000 -
% Equip. Assoc. with S/O 60 60 -
Equip. Assoc. with S/O $ 20,520,000 $ 15,700,000 $ 4,820,000
Other Equipment Cost $ 15,680,000 $ 15,680,000 $ -
Installed Cost, $MM $ 36,200,000 $ 31,380,000 $ 4,820,000
Unit Cost, $/BBL 1,290 1,120 170

Table 3
UOP/FW SDA Advantage
Operating Costs
Option: 1 2 Delta
Case: High Solvent to Low Solvent to -
Oil Ratio Oil Ratio
Solvent-to-oil Ratio 8 5 3
Solvent C4 Mix C4 Mix -
Average Utility Consumptions
(per barrel of feed)
Fuel, MMBTU 0.075 0.056 0.019
Steam, lbs 12.0 10.5 1.5
Power, kW 2.67 1.77 0.90
Incremental Cost
($Bbl Feed)
Fuel 0.293 0.206 0.087
Steam 0.00006 0.00005 0.000
Power 0.179 0.119 0.060
Total 0.342 0.248 0.094
Yearly Cost (28,000 BPD) $3,600,000 $2,500,000 $1,100,000
Unit Cost Reference
Fuel, MMBTU $3.82
3
Steam, 10 lbs $4.83
Power, $/kW $0.067

IDTC Conference
London, England Page 4 of 11
February 2006
WHEN SOLVENT DEASPHALTING IS THE MOST
APPROPRIATE TECHNOLOGY FOR UPGRADING RESIDUE
Edward J. Houde - UOP LLC
Michael J. McGrath - Foster Wheeler USA Corporation

Process Description
The recovered solvent streams from the DAO and
In non-lube oil production applications, regardless pitch recovery sections are cooled/condensed and
of whether a two-product or three-product recycled back to the extraction section for reuse.
configuration is employed, the unit’s design would
Figure 2: Supercritical SDA Process
employ structured packing, supercritical solvent
recovery and the flexibility to utilize a range of
solvent types to achieve the desired separation.

In lube oil production applications, the preferred

extraction device is the multi-stage rotating disk
contactor although at high capacities structured Vacuum
Residue Extractor DAO
packing might also be employed. The rotating Charge Separator

disk contactor, coupled with supercritical solvent

recovery and the flexibility to utilize a range of
solvent types, can achieve the desired separation Pitch
Stripper

while producing a superior quality product. DAO

Stripper

The following brief description of the process flow

scheme is based on a two-product configuration. Pitch DAO

A process schematic (Figure 2) is provided for Optional Resin Recovery (Three-Product

reference. The UOP/FW SDA Unit consists of Configuration)
essentially three major process operations
(extraction, DAO/solvent and asphalt recovery) As the DAO/solvent mix is heated, oils become
which occur within the unit. less soluble in the mix. Initially these oils are the
feedstock’s heavier resin fraction but eventually
The extraction section consists of the extraction become the lighter DAO components. This
vessel, which will typically contain structured provides an opportunity to produce more than one
packing with proprietary internals. In the extractor, DAO product and add flexibility in processing
pre-diluted feed will be contacted with a counter- options. If the option to recover an intermediate
current flow of solvent. Deaphalted oils and resins resin stream is attractive, a resin settler may be
are selectively recovered overhead with the bulk of added between the extractor and DAO separator.
the solvent, and the asphaltenes and more polar This settler can be included in the unit’s initial
resins are rejected (along with a small amount of design or can be added at a later date. In this
solvent) in the bottoms stream. three-product configuration, the DAO/solvent
mixture exiting the extractor vessel is directed to
The extractor’s overhead DAO/solvent mix flows to the resin settler. There, the mixture’s temperature
the DAO separator and solvent recovery section is raised to the point at which the desired
where its temperature is raised above the intermediate-quality resin stream would be sepa-
solvent’s critical conditions. At these conditions rated.
the DAO is no longer soluble in the solvent, and
separates from the solvent by gravity. This This material would then be heated, typically by
section includes heat exchange, a final DAO mix hot oil exchange, before flowing to the resin
heater (steam, hot oil, or direct fired heater), a stripper for solvent recovery.
DAO separator containing proprietary internals,
and a DAO product steam stripper.
The pitch phase, containing some of the solvent, Effect of Process Variables
flows to the pitch recovery section where it is
heated, flashed and steam stripped to remove any The proper operation of the SDA process is
remaining solvent from the pitch product. The affected by several process variables, including:
pitch and solvent recovery section consist of a
pitch mix heater (hot oil or direct fired heater) and
a pitch product steam stripper.

IDTC Conference
London, England Page 5 of 11
February 2006
WHEN SOLVENT DEASPHALTING IS THE MOST
APPROPRIATE TECHNOLOGY FOR UPGRADING RESIDUE
Edward J. Houde - UOP LLC
Michael J. McGrath - Foster Wheeler USA Corporation

Solvent Selection of solvent present in the unit. Consequently, when

frequent feedstock changes might occur, it is
The yield and quality of the products, which are generally more practical to operate the SDA unit
recovered in a SDA unit, are directly related to the with a constant solvent composition and adjust the
solvent composition. As the molecular weight of operation of the extractor. This assumes,
the solvent increases, the yield of DAO also however, that the new feedstock would not be
increases. At the same time, however, the quality processed continuously for an extended period of
of the DAO declines slightly. Because the DAO is time, in which case adjustments to the solvent
usually further processed in a conversion unit composition may be warranted to provide optimum
designed to utilize highly active, metals-sensitive product recovery and product quality.
catalysts which are incapable of economically
processing feedstocks containing more than a few Extraction Temperature
parts-per-million of organometallics, proper solvent
selection must consider both the desired quantity During normal operation, when both the solvent
and quality of the recovered products. composition and the extraction pressure are fixed,
the yields and qualities of the various products
Solvents normally used in the SDA process recovered in the SDA unit are controlled by
include single components such as propane, adjusting the extractor’s operating temperature.
butanes and pentanes, as well as mixtures of For a given solvent composition, Table 4
these components. In most cases the solvent is summarizes the affect of extraction temperature
supplied from LPG products within the refinery and on both the SDA yields and product quality.
at times includes the corresponding olefins. Since
these materials are typically available within a Increasing the extraction temperature reduces the
refinery, their use as SDA solvents is relatively solubility of the heavier components of the
inexpensive. In addition, because the majority of feedstock, improving DAO quality but reducing
the solvent is recirculated within the unit, solvent DAO yield. Subsequent increases in the extractor
makeup rates are small. temperature can further improve the quality of the
DAO by causing additional rejection of the
Propane, being the most selective of the solvents asphaltenic components.
normally considered for a SDA unit, is specified
when the highest quality of recovered DAO is Table 4
required. This typically results in relatively low Affect of extraction temperature on
yields of DAO. Butanes are used either SDA yield and product quality
individually or as a mixture, as SDA solvents when
Extractor Temperature
a reasonably high yield of high-quality DAO is
desired. Finally, pentanes are used when the Low High
maximum yield of DAO is desired. When DAO
compared to the quality of a DAO recovered with a Yield High Low
butane solvent, the DAO recovered from a Quality Low High
pentane solvent would be heavier, would contain Contaminants High Low
more organometallics and Conradson carbon, and
would require more severe downstream Extractor Temperature
processing. This is due to the presence of the Low High
heavier, more-resinous components of the Pitch
feedstock, as these materials generally have Yield Low High
relatively high contaminant contents. Penetration Low High
Softening Point High Low
The SDA unit design typically includes the
capability to operate within a range of solvent There is, however, an upper limit to the extraction
compositions. During actual operation of the unit, temperature which may be used. As the
however, it may be difficult to quickly adjust extraction temperature approaches the solvent’s
solvent composition to offset a change in critical temperature, the DAO rapidly becomes
feedstock quality. This is because of the quantity less and less soluble in the solvent. Because it is

IDTC Conference
London, England Page 6 of 11
February 2006
WHEN SOLVENT DEASPHALTING IS THE MOST
APPROPRIATE TECHNOLOGY FOR UPGRADING RESIDUE
Edward J. Houde - UOP LLC
Michael J. McGrath - Foster Wheeler USA Corporation

difficult to maintain a stable operation at these extracts the more paraffinic components from
conditions, the extractor temperature is typically vacuum residues while rejecting the condensed
maintained below the solvent’s critical ring aromatics. The deasphalted oil product
temperature. quality is characterized by:

Extraction Pressure ∗∗ Higher paraffinicity – The solvent

deasphalted oils show viscosity indices 20 to
The effect of operating pressure is opposite to 40 points higher than the corresponding
temperature but to a lesser extent. In general, the distilled overhead cylinder oils.
higher the operating pressure the more DAO is
extracted at a specific temperature. The operating ∗∗ Lower carbon/hydrogen ratios – The data
pressure of the extractor is based on the indicate carbon/hydrogen ratios of seven, or
composition of the solvent which is being used. lower, are obtainable because of the high
As indicated previously, sufficient operating rejection of condensed ring aromatics to the
pressure must be maintained to ensure the pitch.
solvent/residue mixture in the extractor is in the
liquid state. Although the unit may be designed for ∗∗ Reduced sulfur and nitrogen content - Pilot
a range of operating pressure, once it is in plant and commercial data indicate
operation the extractor pressure is typically not concentrations of these components in the
considered a control variable. DAO are almost always lower than in the
corresponding feed. Blended feeds could
Solvent Recirculation Rate be an exception.

The quantity of solvent contained in the ∗∗ Reduced metal content - DAO with as little
solvent/residue mixture that is charged to the as 1 wppm nickel plus vanadium content
extractor vessel has an impact on extraction has been produced even from Venezuelan
efficiency. As shown in Figure 1, increasing the residues containing 700 to 1,000 wppm of
amount of solvent in the extractor while these metals.
maintaining a constant DAO yield improves the
degree of separation of the individual components ∗∗ Reduced carbon residue - The carbon
and results in the recovery of a higher-quality residue in deasphalted oils is significantly
DAO. lower than for distilled oils of equivalent
viscosity or mid-boiling point.
However, since the quantity of solvent which is
recirculated within the unit is significantly greater The deasphalted oil product yield-quality relation-
than the amount of feedstock being processed, ships obtained when solvent deasphalting typical
any improvement in product quality which results vacuum residues are shown in Figure 3. This data
from an increased solvent recirculation rate must is based on UOP and Foster Wheeler’s extensive
be balanced against the additional operating costs library of pilot plant and commercial solvent
associated with the increased solvent recirculation deasphalting data.
and solvent recovery requirements, and the
increased capital costs associated with the larger Figure 3: SDA Quality Selectivity
equipment sizes. Once the required solvent-to-oil
ratio is established, however, it is usually not
adjusted unless the feed rate is increased and the 100
solvent circulation becomes the limitation on unit
Appearing in Deasphalted Oil, %

90
Sulfur, Nitrogen and Metals

capacity or if there is some other major feedstock 80

change. 70
60
50
lf ur
40 Su
Yield and Product Quality R
30 CC el
ck m
The solvent deasphalting process selectively
20
Ni iu
10 nad
0 Va
0 10 20 30 40 50 60 70 80 90 100
IDTC Conference Deasphalted Oil Yield, Vol-%
London, England Page 7 of 11
February 2006
WHEN SOLVENT DEASPHALTING IS THE MOST
APPROPRIATE TECHNOLOGY FOR UPGRADING RESIDUE
Edward J. Houde - UOP LLC
Michael J. McGrath - Foster Wheeler USA Corporation

Utility Requirements Evaluation of the unit’s many heat integration

options, which is typically done during the basic
For many processing units it is possible to develop design phase may improve utility and capital
utility consumptions based on either the feed requirements. Table 6 illustrates the impact of
capacity or the amount of products produced. several of these options on the unit’s utility
This is not a realistic approach for solvent requirements, based on a two-product UOP/FW
deasphalting since most of the utility SDA unit operating at 100% of design capacity
consumptions are related to the circulation and (28,200 BPSD) at normal operating conditions.
recovery of the solvent. Consequently, solvent
deasphalting utility consumptions depend more The base case for this comparison minimizes heat
upon the solvent circulation rate than upon the integration and takes a conservative approach to
deasphalted oil yield or the unit feed capacity. An equipment sizing, heat recovery, and capital costs.
accurate estimate of solvent deasphalting utility The other options include:
consumptions requires definition of all three of
these parameters. ∗∗ Option 1 - Heat exchange of the DAO product
with the pitch stripper feed. This option
reduces fired fuel requirements by
Summarized below are typical continuous utility
approximately 11.9 MM BTU/hr;
requirements for solvent deasphalting units using
supercritical solvent recovery as described above. ∗∗ Option 2 - Heat exchange of the pitch product
The lube oil case is a low DAO yield, relatively with the DAO phase. This option reduces the
high solvent to oil ratio operation using steam as DAO phase steam heater requirements by
the primary process heat source. The cracking about 18,500 lbs/hr;
stock case is a high DAO yield, low solvent to oil ∗∗ Option 3 - Designing the DAO/solvent mix to
ratio operation using fired heat as the major solvent exchanger for a specific temperature
process heat source. approach. The base case assumes a 5-shell
Table 5 arrangement and a 20°F temperature
approach. Reducing the temperature
Solvent Deasphalting
Typical Utility Requirements, Per Barrel of Feed approach to say 11°F would result in a similar
energy saving as in Option 2. However, the
Lube Oil Cracking
Stock number of shells would increase anywhere
from 2 to 3 times that of the base case.
Fuel Liberated, Btu (LHV) 81,000 55,500
Power KWH 1.5 1.77
Steam (150 psig), lbs 116 10.5
Cooling Water (25ºF rise), 15 Nil
gal

TABLE 6
ESTIMATED UTILITY REQUIREMENTS

Utility Base Option 1 Option 2

Fuel Fired, MMBTU/hr 35.2 23.3 35.2
Electric Power, kW 1192 1163 1192
Steam, x 103 lbs/hr HP (consumed) 78.5 78.5 60.0
Cooling Water, gpm 925 925 360

IDTC Conference
London, England Page 8 of 11
February 2006
WHEN SOLVENT DEASPHALTING IS THE MOST
APPROPRIATE TECHNOLOGY FOR UPGRADING RESIDUE
Edward J. Houde - UOP LLC
Michael J. McGrath - Foster Wheeler USA Corporation

Application of Solvent Deasphalting production of road bitumen or other specialty

asphalts. For instances where a delayed coker is
In determining the best application of solvent available, the pitch may be sent to the coker for
deasphalting the process should be thought of as final conversion and recovery of the remaining oil
a physical separation process analogous in some value. Pitch has also been used commercially as
respects to vacuum distillation, but not limited by feedstock for IGCC and hydrogen production.
the volatility of its products and selective to certain
molecular types. Example of a Recent Commercial Application

It is analogous to vacuum distillation in that it only A recent project that has started up in the Far East
provides separation of products not conversion of is a good example of the applicability of solvent
products. If you change to a lower quality feed it deasphalting in an existing refinery. In this case,
will result in a lower quality of DAO or a lower yield the refiner desired to recover additional cracking
or both. It is not analogous in that it separates by stock from a residual fuel stream. The solvent
molecular types rather than volatility. The result in deasphalting unit was designed to process 6,000
general is that higher boiling components are BPSD (approximately 40,000 kg/hr) of a Middle
recovered in the unit’s pitch product and the lower Eastern crude blend to recover a 50% yield of
boiling components are recovered in the unit’s DAO. UOP/FW/Sulzer’s latest state-of-the-art,
DAO product. Due to solvent selectivity, however, proprietary internals and structured packing were
the pitch will contain low-boiling, highly aromatic used in both the unit’s extractor and DAO
components while the DAO will contain high- separator.
boiling, paraffinic components. For a fuels type
solvent deasphalter, the pitch production will be The addition of the solvent deasphalting unit
minimized for a specific cracking stock (DAO plus allowed the refiner to increase the amount of
VGO) quality by maximizing the lift in the vacuum transportation fuels produced from the refinery.
unit to the limit of VGO quality. This also The recovered pitch, along with some clarified
minimizes the size of the solvent deasphalter. slurry oil, (≤ 20 Liq Vol % of the blend) was used
as a high viscosity residual fuel by an existing
Solvent deasphalting is less expensive to build nearby IPP to produce power. In order to
and operate than other residue conversion compensate for the higher viscosity of the IPP’s
processes, although as noted earlier, it does not fuel blend, for this application, the fuel system and
provide any actual conversion. Therefore, it is burner temperatures were higher than from a tyical
most applicable to recovering the large quantity of solvent deasphalting operation.
high quality oils in light residues while rejecting the
small quantity of asphaltenes and impurities such
UOP/FW Technical Service
as metals and those components that contribute to
carbon residue. In addition, unlike residue
conversion units which benefit from economy of In addition to the benefits provided by UOP and
scale, solvent deasphalting can be economically FW’s broad commercial experience bases,
applied at very small feed rates. licensees have full access to the widest range of
Finally, solvent deasphalting has good applicability support resources available to the refining
where the demand for a low-value residual fuel is industry. This allows UOP/FW to offer licensees
significantly smaller than the production of residual high-quality support from initial conception,
fuel oil based on vacuum tower bottoms through unit design and construction, and
production. continuing on through unit start-up and monitoring
of the operating unit’s performance. These
Utilization of Pitch services include:
∗∗ Pilot plant testing to establish design basis
The utilization of SDA pitch is very much
∗∗ Technical services during unit checkout and
dependant on the local market. The primary outlet
start-up
for the pitch is as fuel, mainly as a blend
component in the residual oil market. Another ∗∗ Ongoing operations monitoring
significant market is as a blend component in the ∗∗ Engineer and operator training programs

IDTC Conference
London, England Page 9 of 11
February 2006
WHEN SOLVENT DEASPHALTING IS THE MOST
APPROPRIATE TECHNOLOGY FOR UPGRADING RESIDUE
Edward J. Houde - UOP LLC
Michael J. McGrath - Foster Wheeler USA Corporation

∗∗ Review of contractor and vendor information

∗∗ Cost estimating and procurement Foster Wheeler and UOP have extensive pilot
plant experience processing a wide range of
∗∗ Metallurgical consultation and equipment feedstocks using solvents ranging from propane
inspection through hexane. This includes deep deasphalting
∗∗ Analytical testing services of vacuum residues derived from Heavy Arabian
∗∗ Modular extractor design (Safaniya), Kuwait, Tia Juana, Heavy Iranian
(Gach Saran), Light Iranian (Agha-Jari, Sassan
UOP/FW have a staff of field operating engineers and Ahwaz), Delta, Heavy Canadian and
and instrument engineers available to provide Canadian crudes, as well as tar sand oil.
technical assistance during new unit
commissioning and start-up. These engineers can FW has also accumulated more than 2,500 hours
also provide immediate technical assistance to the of operation producing pitch above 250°F
licensee if a problem arises on any operating softening point and has produced 400°F capillary
UOP/FW SDA unit. Technology specialists are melting point pitch (which is estimated to
available to our UOP/FW SDA licensees to correspond to 500°F ring & ball softening point).
monitor unit performance and offer advice on unit In addition, Foster Wheeler has run, in the
operations. In addition, licensees can choose laboratory, a water-cooled belt to confirm the
from a wide variety of training courses, including feasibility of flaking high melting point pitch directly
such topics as operations, engineering and from the pitch stripper. This type of belt is used
laboratory and maintenance procedures. commercially to solidify coal tar pitch and could
find successful commercial application with high
Solvent Deasphalting Pilot Plant melting point pitch.

UOP and FW’s extensive experience with various

pilot plant configurations has confirmed their ability
to produce deasphalted oil and pitch products
having physical properties comparable to those
produced in commercial units operating at similar
conditions. An example of this is illustrated in
Table 7, where pilot plant data and commercial
operating results for deasphalting Kuwait vacuum
residue are compared.

IDTC Conference
London, England Page 10 of 11
February 2006
WHEN SOLVENT DEASPHALTING IS THE MOST
APPROPRIATE TECHNOLOGY FOR UPGRADING RESIDUE
Edward J. Houde - UOP LLC
Michael J. McGrath - Foster Wheeler USA Corporation

TABLE 7
SOLVENT DEASPHALTING
PILOT PLANT vs. COMMERCIAL RDC OPERATION

Feedstock: Kuwait Vacuum Residue (1020ºF TBP Cut Point)

Commercial
Pilot Plant Operation

Feed Inspections

Gravity, ºAPI 6.6 7.2

Viscosity @ 210ºF, SSU 7,129 7,229

Con. Carbon Res., % Wt. 20.1 19.2

Sulfur, % Wt. 4.92 5.14

Nickel + Vanadium, PPM 129 --

Deasphalted Oil Yield, % Vol. 32.0 31.0

Deasphalted Oil Inspections

Gravity, ºAPI 19.8 19.5

Viscosity @ 210ºF, SSU 185 194

Con. Carbon Res., % Wt. 1.67 1.64

Sulfur, % Wt. 2.61 2.88

Nickel + Vanadium, PPM 1.0 1.0

Pitch Inspections

Specific Gravity @ 60/60ºF 1.07 1.06

Penetration @ 77ºF 8 10

Softening Point, ºF, R&B 155 151

IDTC Conference
London, England Page 11 of 11
February 2006
Annual Meeting
March 19-21, 2006
Grand America Hotel
Salt Lake City, UT

AM-06-10 Concepts for an Overall Refinery Energy

Solution Through Novel Integration Of FCC
Flue Gas Power Recovery

Presented By:

Keith Couch
Manager, FCC Process
Development
UOP LLC
Des Plaines, IL

Leonard Bell
Process Sales Specialist
UOP LLC
Des Plaines, IL

John Yarborough
Mechanical Specialist
UOP LLC
Des Plaines, IL

National Petrochemical & Refiners Association 1899 L Street, NW 202.457.0480 voice

Suite 1000 202.429.7726 fax
Washington, DC www.npra.org
20036.3896
This paper has been reproduced for the author or authors as a courtesy by the National
Petrochemical & Refiners Association. Publication of this paper does not signify that the
contents necessarily reflect the opinions of the NPRA, its officers, directors, members, or staff.
Requests for authorization to quote or use the contents should be addressed directly to the
author(s)
CONCEPTS FOR AN OVERALL REFINERY ENERGY
SOLUTION THROUGH NOVEL INTEGRATION OF
FCC FLUE GAS POWER RECOVERY

Keith A. Couch and Leonard E. Bell

UOP LLC
Des Plaines, Illinois, USA

INTRODUCTION
Refiners are more focused today than ever on improving utility consumption and reducing stack
emissions. One area receiving significant interest is power recovery from the FCC flue gas,
especially since this power is “clean” in that no additional CO2 is produced or emitted.

While much work has been done over the past 40 years to improve the reliability and operability
of FCC flue gas power recovery systems, the process has remained largely unchanged; that is,
until now. Traditionally, the FCC flue gas power recovery system has all too often been treated
as an “accessory”, tacked on only to higher capacity, higher pressure FCC units in areas of high
electrical cost. In order to make this technology useful for a wider range of FCC operators, UOP
has developed some innovative improvements to the way power recovery systems are
incorporated into the FCC unit. These innovations significantly reduce the capital cost per unit
of energy recovered from FCC unit flue gas in an environmentally friendly manner. These
innovations can potentially double the ROI for a power recovery system when compared to
traditional installations. This has greatly increased the application range of power recovery
systems to FCC capacities for which it was previously considered uneconomical.

AM-06-10
© 2006 UOP LLC. All rights reserved.
Page 1
In this paper UOP will discuss the history of FCC flue gas power recovery, show the economics
associated with implementation of a traditional power recovery system, discuss some of the
recent advancements in process design and the impact they can have on power recovery
economics.

HISTORY OF FCC FLUE GAS POWER RECOVERY

Energy in FCC flue gas has traditionally been recovered in the form of steam generation in a flue
gas cooler. While the application of flue gas coolers is an efficient system for recovering heat
energy, it ignores the potential energy recovery associated with the flue gas rate and pressure. In
1950, a turbo-expander was installed in the flue gas line of an FCC unit in an attempt to convert
the flue gas pressure into shaft power to supplement the utility requirement of the main air
blower. The initial results were extremely poor. After only 750 hours of operation, catalyst fines
in the flue gas had substantially eroded away the turbine blades and casing.

While there are a few exceptions, typical FCC regenerator designs include two stages of
cyclones inside the regenerator. It was originally believed that this efficiency was high enough to
protect the expander from excessive erosion. This proved not to be the case.

In 1963, Shell Oil solved the flue gas fines problem by placing an additional catalyst separator
outside the regenerator at Norco, Louisiana and Oakmont, Canada refineries. This additional
stage of catalyst separation became known as a Third Stage Separator (TSS), and FCC flue gas
power recovery became a sustainable reality.

The regenerator cyclones substantially reduce both the catalyst loading and the particle size
distribution in the flue gas. This allows the TSS cyclone elements to be designed at much higher
velocities, thus higher efficiency, without a concern for erosion of the TSS. Depending on
regenerator design and catalyst systems, a modern TSS is capable of reducing the catalyst in the
flue gas to less than 1 wt-% catalyst larger than 10 microns. This is more than sufficient to
provide long-term reliable protection of the power recovery expander.

There are several possible Power Recovery System configurations that can be incorporated into
both new and existing FCC Unit designs. Selection of the specific equipment type and
arrangement is always refinery specific, and is generally a balance of utility requirements,
process optimization, and preference between different configurations.

Amongst the various configurations, there are basically two types of power recovery applications
to consider: 1) a four or five body power recovery train in which the expander is directly coupled
to the main air blower to provide direct shaft power, and 2) a two, three or four body power

AM-06-10
Page 2
recovery train in which the expander is coupled to a generator to produce electrical power. These
configurations will be discussed in turn.

TRADITIONAL FCC FLUE GAS POWER RECOVERY SYSTEMS

Four and Five-Body Coupled Trains

Between 1963 and 1981, 18 power recovery applications were commissioned industry-wide.
These were typically five-body trains; consisting of a hot gas expander, main air blower, steam
turbine, motor/generator and gear box as necessary. A five-body train is shown in Figure 1. This
configuration was historically the most common power recovery system for new unit
installations.
FIGURE 1
Traditional 5-Body Power Recovery Train

Flue Air Air to

Gas In Regenerator
Exhaust
Inlet Steam Gear Motor/
Guide Turbine Box Generator
Flue Vanes
Gas
Inlet
Main Air Electrical
Blower Connection to
Power Grid
Exhaust
Steam
Expander
Outlet
Steam
Inlet

In this configuration, the expander is coupled to the main air blower and provides a direct
transfer of energy to the shaft. The direct transfer of energy to the main air blower minimizes
power transfer losses, and is the most energy efficient configuration.

The steam turbine is used to start up the train. The shaft speed is increased to match the electrical
frequency of the motor (the “synchronous speed”) to that of the power grid and the electrical
breaker is closed. Once the breaker is closed, the shaft speed of the train is fixed to the frequency
of the power grid. The air flow rate is controlled by adjusting the inlet guide vanes. Generally,
the combination of a steam turbine and motor can provide the required power to operate the air
blower at design conditions, with the expander out of service. After enough flue gas is present,
and process conditions are stable, the expander can be commissioned. An overview schematic of
the traditional power recovery train is shown in Figure 2.
AM-06-10
Page 3
FIGURE 2
Traditional Power Recovery System – “Five-Body Train”

The motor/generator imports or exports power as required to maintain a constant train shaft
speed. If more energy is supplied by the expander than is required by the blower there is a
surplus of electricity generated and exported to the power grid. Conversely, if the blower power
requirement is higher than the expander is providing, electricity is consumed by the motor to
maintain the train at normal speed. With the expander coupled to the main air blower, in the
event of a breaker disconnect the shaft power requirement of the MAB acts as over-speed
protection for the expander.

With a power recovery system, butterfly valves are used in the flue gas line to control the
differential pressure between the reactor and regenerator. A dedicated, high speed “power
recovery control system” performs all process control functions. Fundamentally, the regenerator-
reactor pressure differential controller (PDIC) adjusts the expander inlet valve to control the
differential pressure, and only opens the bypass valve when the expander is out of service or
expander maximum throughput is reached. This control scheme maximizes the power generation
potential by minimizing the amount of flue gas that is diverted around the expander.

AM-06-10
Page 4
With a traditional five-body power recovery train (PRT) if the expander requires any repair or
maintenance, the entire FCC Unit has to be shut down. The business interruption costs associated
with shutting down the entire FCC unit can be substantial, and rapidly negate the economic
advantage associated with the power recovery system. With these concerns in mind, there was a
desire by many refiners to decouple the power recovery expander from the main air blower shaft.
This was initially a very challenging problem that was solved in the early 1980s, and ushered in
the Gen Set power recovery system.

Gen Set Power Recovery Train

The gen set PRT is a “stand-alone” system in which the expander is connected to a generator and
the main air blower is installed as a separate machine. An overview schematic of a modern Gen
Set power recovery system is shown in Figure 3. By removing the main air blower from the
power recovery train, the shaft load associated with the blower is eliminated. The main concern
with this configuration was that in the event of a breaker-disconnect the shaft load drops
essentially to zero and the expander could speed up in an uncontrolled manner, resulting in a
potential over-speed. These concerns led to the development of fast acting control valves, high
speed electro-hydraulic actuators, and improved instrumentation and control systems that rapidly
divert flue gas out of the expander to decelerate the train. In 1983, one of the first gen set power
recovery systems was commissioned at Saras, S.p.A., in Sardinia, Italy.

AM-06-10
Page 5
FIGURE 3
Gen Set Power Recovery System

Split
HSS Range Flue
> Gas
Orifice
Chamber
PIC PDIC

Butterfly
Third
Valves
Stage
Separator Flue Gas Electrostatic
Combustor Cooler Precipitator
Style
Regenerator Critical
Flow
Nozzle

Gear
Box Generator

To / From
Reactor
Expander
Gear Motor
Box

Main Air
Blower

With a modern Gen Set power recovery application, FCC unit down time associated with either
maintenance or failure of the power recovery expander is essentially eliminated. Operation of the
FCC unit can be maintained while the PRT is isolated to complete any repair or routine
maintenance work.

The decision to couple the power recovery expander to the MAB shaft or install a Gen Set PRT
is a question that each refiner must address. With the current cost of FCC unit downtime, more
refiners are seeing the potential economic advantage of the Gen Set power recovery system.

COMMERCIAL HISTORY OF POWER RECOVERY SYSTEMS

Through the 1970s and 80s refinery investment in FCC flue gas power recovery tracked very
closely with electrical prices. When compared on a 2004 constant dollar basis, electrical prices
hit a peak across 1981-86. By this time, power recovery had achieved broad based acceptance
across the industry, and the technology experienced rapid growth. See Figure 4. The numbers of
“New PRT Installations” (107 total) are based on the equipment commissioning date. It should
be noted that additional power recovery trains have likely been installed in China and the FSU
that were not publicly disclosed and are not represented herein.

AM-06-10
Page 6
FIGURE 4
Historical Power Costs & PRT Installations
12.00 10
Total New PRT Instillations
Coal Price, x10 $/short ton
9
Natural Gas Price, Wellhead ($ / k Cubic Feet)
10.00 Average Retail Price of Electricity, Industrial (Cents/KWH)
8

# of New PRT Instillations

7
Price, Constant $ -2004

8.00
6

6.00 5

4
4.00
3

2
2.00
1

0.00 0
1973 1978 1983 1988 1993 1998 2003

* Historical pricing data for Figure 7 provided by U.S. Department of Energy

Although natural gas prices have been rising sharply in recent years, since about 1993 coal has
dramatically outpaced all other fuels for electrical power generation. Both the availability and
low price of coal has maintained electrical rates at historic lows since 1999, but are only recently
starting to increase as shown in Figure 4.

In contrast with the current low electrical rates, UOP has observed a recent resurgence in the
industry interest in FCC flue gas power recovery systems. This interest appears to be primarily
driven by refiners focusing on direct returns by lowering their operating cost, and indirect returns
on investor confidence by improving the Energy Intensity Index (EII) of their operations and
reducing environmental emissions.

Energy Intensity Index

The EII was developed by Solomon Associates in 1981 to compare energy consumption among
fuels refineries. Standard energy consumption per barrel of utilized capacity was established for
each process and process types. The factor for fluid catalytic cracking units is based on an FCC
configuration without a flue gas power recovery expander. The EII is essentially the ratio of the
actual energy consumption divided by the sum of the refinery standard energy consumption. The
installation of a flue gas power recovery system reduces the actual energy purchased. As such,
the EII for refiners that have an FCC unit flue gas power recovery expander is lower than the
standard factor, representing better than standard energy efficiency.

AM-06-10
Page 7
Environmental Emissions
The application of an FCC flue gas power recovery system is “green” with respect to electrical
power generation. No additional CO2, SOx or NOx are created in association with the power
generated. This can provide both permitting and economic benefits to the refiner. The economic
benefit is going to be discussed later in this paper.

The impact on the refiner operating with a lower EII and good neighbor emissions stewardship
improves investor confidence that the refinery is being properly managed. As a result, many
refiners will accept lower returns on capital projects focused on energy optimization. However,
the projects still have to be economically attractive. With the recent low electrical rates this has
proved particularly challenging for traditional FCC flue gas power recovery and has focused
UOP on inventing new ways to improve the economic viability for power recovery systems and
improving the reliability of the system.

With renewed interest in FCC flue gas power recovery, UOP reviewed many of the historical
designs and realized that very little had changed since the early 1970s. Opportunities arising
from recent advancements in component equipment designs (expander, actuators, instruments,
and TSS) had not been utilized.

As emission requirements have tightened, particulate matter in the FCC flue gas has been
controlled by the installation of an electrostatic precipitator, wet gas scrubber or barrier filter.
Each of these emission control devices requires considerable capital and operating expense, but
has no financial Return on Investment (ROI) beyond supporting a permit to operate.

UOP ADVANCEMENTS IN FCC FLUE GAS POWER RECOVERY SYSTEMS

Over the past three years, UOP has developed innovative improvements in the design and
application of power recovery systems. These innovations have been targeted to optimize the
efficiency of energy recovered from the FCC flue gas stream in an environmentally friendly
manner and at a significantly lower total erected cost. With these innovations, project economics
have been significantly improved when compared to traditional designs, expanding the
application range of flue gas power recovery to lower capacity units, for which power recovery
systems were previously considered uneconomical.

The remainder of this paper will discuss some of UOP’s recent FCC flue gas power recovery
innovations. With each innovation a supportive economic case study will be shown that
progresses one advancement to the next, starting with the traditional maximum electrical power
generation configuration. Although the economic analyses are cited as case studies, they are
closely based on recent commercial projects.
AM-06-10
Page 8
Flue Gas Basis
To start this series of case studies, a relatively common flue gas basis was chosen which meets a
“middle of the road” FCC operation. The following conditions were used.

Feed rate: 50,000 BPSD

Coke yield: 5.3 wt-%FF
Flue gas excess oxygen: 1.0 mol-%
Operating days per year: 365
Flue gas rate: 533,000 lb/hr
Flue gas temperature: 1350°F
Regenerator pressure: 34.0 psig

Utility Costs
The utility costs used for the case studies were based on 2005 U.S. Gulf Coast costs.

Erected Cost Analysis

The unit costs highlighted for each power recovery installation case study presented in this paper
were developed by UOP using historical data generated from recent similar power recovery
installation projects, to support each associated economic analysis. These costs have been
developed on a U.S. Gulf Coast installation basis, and include recent vendor quotations for most
major equipment items.

These costs are represented as total erected costs, and as such, exclude all project costs (licensor
basic design, royalties, and T&K fees, spare parts, start-up services, training, and owners costs),
as well as project contingencies, as these costs are unit-specific and carry a wide range of
variability.

Nevertheless, every attempt has been made to develop each of these case specific erected costs
on a consistent design and installation basis to ensure comparative accuracy and effectively
support the resultant ROI representations.

CASE STUDY #1: (TRADITIONAL POWER RECOVERY SYSTEMS)

This first case considers a traditional FCC flue gas power recovery system as detailed in Figure
3. In this configuration, electrical power generation is maximized by feeding the highest
temperature, highest pressure flue gas to the inlet of the power recovery expander. Residual heat
in the flue gas down stream of the expander is recovered in the form of HP steam generation.
From this case, 13.78 MW of electrical power are generated. Since the application of the
expander reduces the temperature to the flue gas cooler, there is a loss of HP steam production

AM-06-10
Page 9
that must be made up in the boiler house. A debit has been applied to the economics reflecting a
pound-for-pound shift in HP steam production from the FCC unit to the boiler house.

The economic analysis for the installation of a traditional power recovery system is shown in
Table 1. The ROI for this project is presented for 25, 30, and 35 percent discounted cash flow
factors, and shows a very marginal return between 9.1 – 10.6%. This is an example of why it has
been difficult in recent years to economically justify the installation of a power recovery system
on moderate to small sized units.

TABLE 1
Utility Analysis and ROI - Traditional FCC Power Recovery System

Erected Cost, HP Steam Electrical DCF ROI

(MM-$) (lb/hr) Power, (MW) 25% 30% 35%
Base Case N/A 110,300 N/A N/A N/A N/A
Add Traditional PRT $28,900,000 69,800 13.78 12.7 11.9 11.0

The economics of major capital projects often improve with larger size units due to economies of
scale. This is true for the application of FCC flue gas power recovery systems. Figures 5 & 6
show a comparison of erected costs and ROIs for various capacity units, pegged to the conditions
for the 50,000 BPSD case study of this paper and a scale-up to a 125,000 BPSD unit. The two
lines in each Figure show the trend lines for a traditional power recovery and the new UOP
power recovery system. As the size of the unit is decreased, a minimum erected cost level is
reached, which can result in falling ROI levels with smaller units. Conversely, on larger units,
the incremental increase in erected costs leads to higher ROIs.

The remainder of this paper is going to discuss some of the aspects of the new UOP power
recovery system, and how we have been able to reduce the total erected cost and improve energy
integrations for a step change improvement in ROI over base.

AM-06-10
Page 10
FIGURE 5
Relative Erected Cost Impact with Varying Size FCC Units

Base
Relative Erected Cost

Base
-45%
-30%

Traditional PRT System

New UOP PRT System

0 20,000 40,000 60,000 80,000 100,000 120,000 140,000

Unit Throughput, (BPSD)

FIGURE 6
Relative ROI Impact with Varying Size FCC Units

2.2 x's Base

Return on Investment, (%)

2.0 x's Base

Base
Base
Traditional PRT S ystem
New UOP PRT S ystem

0 20,000 40,000 60,000 80,000 100,000 120,000 140,000

Unit Throughput, (BPSD)

AM-06-10
Page 11
CASE STUDY #2: (TRADITIONAL PRT WITH TSS INTEGRATED EXPANDER
BYPASS LINE)
Advancements in TSS design provided the first power recovery system optimization opportunity
by reducing the capital costs of the TSS, and associated flue gas duct and structure. The new
UOP TSS is about 40 percent smaller than other TSS offerings for the same capacity, making it
less expensive to fabricate, easier to install, and better suited where plot space is a premium.
Figure 7 shows an equal capacity comparison of the older radial flow TSS and the new UOP TSS
design.

The most significant improvement in the design is that the UOP TSS utilizes axial flow for
catalyst/gas separation. The flue gas flow is maintained essentially in one direction - in the top
and out the bottom of the unit. Axial flow distribution minimizes the potential for solids re-
entrainment resulting from changes in directional gas flow and resultant eddy current formations.
More importantly, reducing the directional flow changes minimizes pressure loss across the TSS
by 0.25 to 0.50 psi. This translates directly into more power recovery potential across the
expander.

In addition to being smaller, the clean gas outlet is located on the lower side of the vessel. This
minimizes the amount of steel structure and hot wall flue gas duct (typically 304H stainless steel)
between the TSS and the expander inlet.

FIGURE 7
Third Stage Separator – Equal Capacity Comparison

AM-06-10
Page 12
Proper design of the expander inlet / outlet lines is extremely critical to the overall operation and
reliability of the power recovery expander. The allowable nozzle loadings on the expander inlet
and outlet nozzle are extremely small. Great care must be taken in the detailed engineering process
to ensure that the nozzle loadings are maintained within allowable parameters across the entire
operating range as well as transient conditions of the system.

In older power recovery system designs, the expander bypass line was typically designed as hot-
wall pipe connected directly to the expander inlet line “minimum distance” from the outlet of the
TSS. See Figure 8. The objective of this configuration was to minimize the loading effects that the
bypass line imparts on the expander inlet nozzle during intermittent or transient use. The bypass
line can be in service during all modes of operation; startup, shutdown and normal operation. As
the bypass valve opens, the flow of hot flue gas causes the flue gas duct to heat up and thermally
expand. The resultant duct expansion imposes a great deal of force loading and rotational moment
on the expander inlet line, making the piping system both difficult to design and costly to install.

FIGURE 8
Traditional Expander Bypass Line Installation

Flue Gas from

Regenerator
S
T
Expander Bypass Line A
TSS
TSS C
Flue Gas
Cooler K
Min
Gear
E G
M

TSS Underflow

Where: E = Expander, Gear, G = Generator

AM-06-10
Page 13
To alleviate this problem, UOP has added a second clean gas outlet nozzle to the TSS and attached
the expander bypass line as a lower cost cold-wall line, directly to the TSS vessel. See Figure 9.

FIGURE 9
UOP Modified Expander Bypass Line Installation

Flue Gas from

Regenerator
S
T
Expander Bypass Line A
TSS
TSS C
Flue Gas K
Cooler
Gear
E G

TSS Underflow

Where: TSS = Third Stage Separator, E = Expander, G = Gear, M/G = Motor Generator

This configuration has many significant benefits that increase the overall reliability, operability,
and cost effectiveness of the system. In this design, the TSS acts as a fixed anchor point in the duct
design for both the expander inlet and bypass lines.

The line from the TSS to the inlet of the expander becomes a very clean, minimum impact design,
allowing for shorter duct length between the expander and the TSS. The transient loads applied to
the expander inlet nozzle associated with intermittent or normal use of the bypass line are
minimized, essentially reduced to only the axial pressure thrust with varying gas flows. The bypass
line also becomes much shorter in length and can be of cold wall design. Both of these provide for
lower overall design and installation costs as well as operational and maintenance benefits.

The combined implementation of an axial flow UOP TSS with a side connected clean gas outlet
and integral cold-wall bypass line reduces the detailed piping design requirements, structural
steelwork, large diameter piping, pipe supports, and expansion joints for the system. This lowers
the total erected cost for this case study from $28.9 MM to $27.1 MM, with a corresponding
improvement in ROI as shown in Table 2.

AM-06-10
Page 14
TABLE 2
Utility Analysis and ROI – UOP TSS and TSS Anchored Bypass Line

Erected HP Steam Electrical DCF ROI

Cost, ($) (lb/hr) Power, (MW) 25% 30% 35%
Add Traditional $28,900,000 69,800 13.78 12.7 11.9 11.0
Five-Body PRT
TSS Integral $27,100,000 69,800 13.78 13.6 12.7 11.8
Bypass Line

Although the economic benefits seen here with 1) the addition of a PRT and 2) the installation of a
UOP TSS are relatively small, it leads into a progression of additional technology improvements
and costs reductions that greatly improve the economics of flue gas power recovery.

CASE STUDY #3: (PROCESS INTEGRATED STEAM LETDOWN TURBINE)

The integration of a steam turbine as a supplemental driver to assist with startup of the PRT is
relatively common. In traditional operation, the flue gas flow is maximized to the expander and the
steam flow to the turbine is reduced to the minimum required to keep the turbine in operation. At
this point, the steam turbine becomes a marginally utilized asset until the next FCC shutdown.

Integration of the steam turbine provides a means for the refiner to supplement electrical power
generation from the PRT. However, in designs of the past, the generator has been sized according to
the rated capacity of the expander. As such, using the steam turbine to supplement electrical power
generation has been limited to times when the FCC unit is in turndown operations and not making
the design power output. In considering ways to maximize the use of existing assets, the PRT steam
turbine has been a significantly under-utilized piece of equipment.

Most refineries operate a boiler house that generates a single level of high pressure steam. Lower
levels of steam are supplied by successive letdown stations to the medium- and low-pressure steam
headers. Additional LP steam is also often generated by exhausting steam turbines into the LP
header. A typical steam letdown configuration is shown in Figure 10.

AM-06-10
Page 15
FIGURE 10
Typical Steam Letdown Configuration

When steam is let down across a control valve, the potential energy of the steam is reduced without
work being done. With the addition of a steam turbine in the FCC power recovery train as shown in
Figure 11, there is a way to capture a significant amount of this lost energy.

FIGURE 11
UOP Steam Letdown Configuration
HP Steam Header
(600 or 450# typ)
Boiler
House
FCC
Power Recovery Train
Gear
G
E T

PIC
MP Steam Header
(150# typ)

LP Steam Header
Refinery PIC
(50# typ)
Turbine
Exhaust

Where: E = Expander, Gear, G = Generator

AM-06-10
Page 16
In this configuration, HP or MP steam can be let down efficiently across the turbine into the lower
pressure steam headers. The energy transferred to the generator shaft is used to produce
supplemental electrical power. To optimize this configuration, the PRT motor/generator needs to be
sized to accommodate the proper steam let-down requirements to meet the refinery’s needs. As
depicted, multiple levels of steam letdown can be accommodated through a single turbine. This
utility integration is a process design that extends beyond the battery limits of the FCC unit, and
allows the refiner to optimize the economics of operating their facility-wide steam and electrical
systems.

When installing an FCC flue gas power recovery system, most of the auxiliary equipment is already
required; i.e. the generator, 13.8 kVa cable, switches gear, foundation, electrical controls and
substation. The incremental cost of adding the steam letdown turbine to the power recovery train is
low compared to the potential energy recovered, and can significantly increase the return on
investment for installation of a power recovery system in the FCC unit.

If we focus on the battery limits of the FCC unit, there are additional energy integration
opportunities available than can be used to help maximize use of available assets; in this case, the
PRT steam letdown steam turbine. There are several MP steam flows to the reactor including the
riser lift steam, feed distributor dispersion steam, spent catalyst stripping steam, and reactor
fluidization steam. The steam that is injected into the reactor is heated up to the reactor operating
temperature, and is one of the loads on the overall heat balance. From a process standpoint, we do
not want the additional heat input that the superheat of the MP steam provides. However, LP steam
is simply too low in pressure to be used for these applications.

With the integration of a flue gas power recovery system, the normal FCC process steam can be
routed through a letdown turbine on the PRT as shown in Figure 12. In this manner, the excess
superheat and pressure energy of the steam is transferred as shaft power to the PRT and used to
either supplement the blower power requirement or to produce electricity in the generator.

AM-06-10
Page 17
FIGURE 12
UOP Process Integrated Steam Turbine
Products to
Main Column

HP/MP
Stripping Steam Outlet
Steam

Feed Gear
Distributor
Steam G
E T
Steam
Outlet

PIC
LP Steam

Lift
Steam

PIC
PIC
MP Steam

~100 psig
MP Steam Header Steam Letdown

In application, the higher the steam pressure supplied to the turbine, the greater the economic return
on integrated electrical power generation. The turbine exhaust pressure is a variable with which the
operator can control the amount of superheat remaining in the steam; the lower the exhaust pressure,
the lower the remaining superheat. As more energy is removed, the higher the electrical power
generation from the PRT and the higher the heat load on the reactor. While the temperature of the
steam is a relatively small heat load on the reactor, the lower the steam temperature, the higher the
catalyst-to-oil ratio in the unit for improved product selectivities.

The addition of a steam letdown turbine within the power recovery train increases the total
erected cost of the system from $27.1 MM to $28.4 MM as shown in Table 3. With the steam
conditions used, integrating just the reactor, riser, and stripper steam letdown covers the
additional cost of the turbine and steam integration. If the steam source is changed from MP to
HP steam, the ROI is improved by +1.7 percent.
AM-06-10
Page 18
TABLE 3
Utility Analysis and ROI – Integrated Process Steam Letdown

Erected HP Steam Electrical DCF ROI

Cost, ($) (lb/hr) Power, (MW) 25% 30% 35%
Add Traditional $28,900,000 69,800 13.78 12.7 11.9 11.0
Five-Body PRT
TSS Integrated $27,100,000 69,800 13.78 13.6 12.7 11.8
Bypass Line
Reactor Riser Steam MP $28,400,000 69,800 14.14 13.6 12.7 11.8
Integrated Turbine
Reactor Riser Steam $28,400,000 69,800 15.60 15.5 14.5 13.5
HP Letdown Turbine

With the integration of a steam turbine, multiple levels of letdown can be simultaneously
incorporated into the system. To provide a basis for the economic impact that integrated steam
letdown can have, Table 4 shows the additional power generation and economic improvement
per 10,000 pounds of steam letdown for three different levels of steam.

TABLE 4
Utility Analysis and ROI – Per 10,000 lb/hr Steam Letdown

Steam Letdown Heat Rate Steam Rate Generator DCF ROI

(lb/kWh) (lb/hr) Power, (MW)
600# to 90# Header 25.87 10,000 +0.39 +0.5
600# to 150# Header 33.10 10,000 +0.30 +0.4
150# to 50# Header 116.00 10,000 +0.09 +0.2

For example, if 30,000 lb/hr of steam is let down from the 600# HP to the 150# MP steam header
for a 25 percent DCF bracket, the subsequent increase in ROI and recovered electrical power
would be:
ROI = 13.0 + (3 * 0.5) MW = 15.6 + (3* 0.39)
ROI = 14.5% and, MW = 16.77

AM-06-10
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CASE STUDY #4: TEMPERATURE CONTROLLED EXPANDER INLET
With the traditional approach to power recovery, electrical power generation is maximized by
directing the highest temperature, highest pressure flue gas to the inlet of the expander. The energy
recovered across the expander results in a flue gas temperature reduction of 150 - 400°F and a
minimum exhaust pressure. The flue gas is then routed to a low pressure flue gas cooler for residual
energy recovery in the form of steam production. However, due to the temperature reduction, the
quantity of the steam generation is lower than before the expander was placed into service. Even
with the most efficient cooler designs, installing a power recovery expander upstream of a flue gas
cooler to maximize electrical power generation can result in a 20-30 percent reduction in high
pressure steam production that must be financially off-set with the value of electrical power
generation for the installation of a power recovery system to be economically attractive.

Considering the dynamic balance between steam and electricity costs, UOP evaluated several
options to improve the economics for installing a power recovery system. In a traditionally applied
power recovery system, the operating temperature of the regenerator dictates the inlet
temperature to the expander. The higher flue gas temperature at the expander inlet duct requires
the use of expensive stainless steel duct.

In the new temperature-controlled expander inlet design shown in Figure 13, the expander is
placed downstream of a high pressure flue gas cooler, reducing the metallurgy requirement of the
entire power recovery system to lower cost carbon steel, resulting in a total erected cost that is
potentially 30–40 percent lower than that of a traditional system design, depending on the capacity
of the unit.

This provides the refiner with another means to help optimize the economics of their overall
refinery utility systems between maximum HP steam generation and maximum electrical power
generation at 1050°F maximum expander inlet.

AM-06-10
Page 20
FIGURE 13
UOP Temperature Controlled Power Recovery System

HP Steam BFW

HP
Steam Economizer Stack
Regenerator Gen

Steam Inlet

Gear
Air Inlet TSS
E T G

Blower Motor

In this new system, the flue gas from the regenerator is first routed through a flue gas cooler,
then to the power recovery train. A bypass line is installed around the flue gas cooler to provide a
means to control the inlet temperature to the expander. This allows the refiner to optimize energy
production from the FCC unit between steam and electrical power. The control system is
configured to provide load-following variable peak response control of both the refinery HP steam
system and electrical distribution. If the refinery HP steam requirement drops, or if the refinery is
close to an electrical surcharge threshold, the flue gas cooler bypass line is opened to direct
additional hot flue gas to the expander to produce additional electrical power. This allows the
refiner the capability to optimize the HP steam generation and electrical power generation as
utility economics shift, independent of the operation of the FCC reactor / regenerator.

The lower flue gas temperature downstream of the flue gas cooler allows the entire power recovery
system (vessels, control valves, expansion joints, piping, and duct work) to be designed and
installed with lower cost carbon steel materials as opposed to the higher cost stainless steel and cold
wall refractory lined duct work required by the traditional system. The lower temperature system
design results in less thermal movement of the flue gas duct, reducing the size, type, and quantity of
expansion joints required. This further reduces the erected cost of the system. Even though the
electrical power generated with the lower temperature expander inlet is reduced from that of the
traditional system, as shown in Table 5, the substantially lower cost of the system far exceeds the
reduction in electrical power generation, resulting in a significant step change increase in ROI.
AM-06-10
Page 21
TABLE 5
Utility Analysis and ROI – Temperature Controlled Expander Inlet

Erected HP Steam Electrical DCF ROI

Cost, ($) (lb/hr) Power, (MW) 25% 30% 35%
Base Case N/A 110,300 N/A N/A N/A N/A
Add Traditional $28,900,000 69,800 13.78 12.7 11.9 11.0
Five-Body PRT
TSS Integrated $27,100,000 69,800 13.78 13.6 12.7 11.8
Bypass Line
Reactor Riser HP Steam $28,400,000 69,800 15.60 15.5 14.5 13.5
Letdown Turbine
Temperature Controlled $19,900,000 79,300 12.61 24.2 22.6 21.0
Expander Inlet

In addition to the economics presented in Table 5, the lower inlet temperature to the expander
increases the long term reliability of the system and helps minimize expander blade erosion and
power recovery loss over time. The cooler catalyst particles that pass over the expander blades are
much less apt to fuse into catalyst deposits on the blades and casing, further improving the system
reliability as a function of reducing expander blade tip erosion and tip-rub-induced shaft vibration.

CASE STUDY #5: ENVIRONMENTAL CONSIDERATIONS

The application of an FCC flue gas power recovery system is “green” with respect to electrical
power generation in that no additional CO2, SOx or NOx are created in association with the
power generated. A paper issued by Colin High4 detailed the value of emission reductions as
shown in Table 6:

TABLE 6
Value of Emissions Reductions – $ / metric ton

$ / metric ton
Carbon Dioxide (CO2) 3–5
Sulfur Dioxide (SO2) 300 – 600
Nitrogen Oxide (NOx) 3,000 – 10,000

AM-06-10
Page 22
To assess the value of generating “green” power with a power recovery system, the marginal fuel
for electrical production was considered to be natural gas. Every kW-hr produced in the power
recovery system results in a kW-hr lower requirement from a cogeneration plant. With the use of
natural gas, the SOx reduction is near zero, the CO2 reduction is proportional to the fuel
consumption, and the NOx reduction is based on the use of low NOx burners with an emission of
40 ppm NOx in the flue gas. The value of emission reductions used for the economic analysis is
based on the average of the ranges shown in Table 6. With a fuel gas heat rate of 9,090
BTU/kW-hr, the resultant value for emissions reductions is tabulated in Table 7.

TABLE 7
Emissions Reductions– metric ton/MW-hr

metric ton/MW-hr ¢ / kW-hr*

Carbon Dioxide (CO2) 0.554 0.222
Sulfur Dioxide (SO2) 0.000 0.000
Nitrogen Oxide (NOx) 0.000174 0.113
0.335
*Based on Average Values of Table 8

The economic impact of the emissions reduction in association with installation of a power
recovery system is noticeable, and as shown in Table 8, further improves the ROI for a power
recovery system.
TABLE 8
Utility Analysis and ROI – Emissions Credit

Erected HP Steam Electrical DCF ROI

AM-06-10
Page 23
SUMMARY
Increased global focus on reducing energy consumption and emissions are working together to
make FCC flue gas power recovery more attractive. Even in this environment, the economics
associated with a traditional power recovery system can be marginal for an average sized FCC
Unit. However, UOP has developed a series of novel improvements to the traditional scheme that
make it an attractive investment across a broader range of FCC capacities at the current price of
electricity.

The improvements discussed herein, while novel in application, are all supported by proven
technologies and serve to reduce capital cost or improve efficiency and availability. The ‘TSS
Integrated Bypass Line’ reduces capital cost of the expander inlet line and is made possible by
UOP’s commercially proven new TSS design. The ‘Reactor Riser Steam Letdown Turbine’
utilizes existing turbine technology to improve efficiency. The ‘Temperature Controlled
Expander Inlet’ utilizes existing cooler technology, along with a turbine, to reduce capital cost,
improve efficiency and improve availability.

While an abundant supply of low cost coal has helped keep electricity prices in check, there are
signs that the price of coal is on the rise. An increase in electricity prices to the inflation-
adjusted average of 1973 to 1988, with all else equal, would result in a 70 percent increase in the
ROI for all cases considered. Because UOP believes that the long-term factors that drive energy
efficiency are on the rise and will remain so for the foreseeable future, UOP remains committed
to improving the process behind power recovery from FCC flue gas systems.

The concepts presented in this paper provide a glimpse at some of the recent work UOP has been
performing on FCC flue gas power recovery. New ideas and new opportunities are being
developed that build upon the recent technology advancements. UOP is currently working on a
new power recovery system that further reduces total erected cost and increases the overall
power recovered to further improve ROI. The newer system significantly reduces required plot
space and allows the refiner to potentially meet current and future particulate matter stack
emission requirements.

AM-06-10
Page 24
ACKNOWLEDGEMENTS
The authors of this paper would like to express their thanks to the following companies and
individuals, for their assistance in providing data and/or support that have helped make this paper
a reality.

1. Dresser-Rand Company - David. M Vincent, David M. Hargreaves, Kenneth J. Reading,

Stephen W. Knight, and George H. Seamon –Assistance with equipment costs, power
generation estimates, equipment drawings and historical perspectives on FCC flue gas power
recovery.

2. UOP – FCC Engineering Tech Center – Daniel N. Myers –– Assistance with Flue Gas
Cooler steam estimates, and historical perspectives on FCC flue gas power recovery.

3. UOP - FCC Engineering Tech Center - John Yarborough – Assistance with pipe stress
analyses, large bore piping design, proper equipment layout, and 3-D graphics used in the
presentation.

REFERENCES
1. V. J. Memmott, and B. Dodds, “Innovative Technology Meets Processing and Environmental
Goals: Flying J Commissions New MSCC and TSS”, National Petrochemical & Refiners
Association (NPRA) Annual Meeting, Paper AM-03-13, March 2003

2. K. A. Couch, K. D. Seibert, and P. J. Van Opdorp, “Controlling FCC Yields and emissions –
UOP Technology for a Changing Environment”, National Petrochemical & Refiners
Association (NPRA) Annual Meeting, Paper AM-04-45, March 2004

3. Solomon Associates, “Refinery Comparative Performance Analysis Methodology",

Copyrighted 2001

4. Colin High, “Air Emissions Reductions and Value from Green Power”, Ninth National Green
Power Marketing Conference, Albany, New York, Oct 4-6, 2004

5. K. J. Reading, “Expander Controls Past, Present and Future”, Third International Expanders
Users' Council, Houston, Texas, June 1995

AM-06-10
Page 25
6. K. J. Reading, S. Rubino, and T. Kociuba, “The Application of Large Induction Generator to
a Fluid Catalytic Cracking Power Recovery Train”, IEEE / IAS Annual Meeting, Denver, CO
October, 1986

UOP LLC
25 East Algonquin Road
Des Plaines, IL 60017-5017

Page 26
Annual Meeting
March 19-21, 2006
Grand America Hotel
Salt Lake City, UT

AM-06-16 Changing Refinery Configuration for Heavy

and Synthetic Crude Processing

Presented By:

Gary Brierley
Technology Manager
UOP LLC
Des Plaines, IL

Visnja Gembicki
Marketing Manager,
Refining
UOP LLC
Des Plaines, IL

Tim Cowan
Senior Development
Engineer
UOP LLC
Des Plaines, IL

National Petrochemical & Refiners Association 1899 L Street, NW 202.457.0480 voice

CHANGING REFINERY CONFIGURATION FOR HEAVY

AND SYNTHETIC CRUDE PROCESSING

Gary R. Brierley, Visnja A. Gembicki and Tim M. Cowan

UOP LLC
Des Plaines, Illinois, USA

INTRODUCTION
Reduced availability of light conventional crudes in the future will create demand for new crude
sources that will necessitate refinery configuration changes. The production of heavy crudes,
synthetic crudes, and bitumen blends is growing, and the supply of bitumen-derived crudes is
expected to reach almost three million barrels per day by the year 20151. A plethora of synthetic
crudes and bitumen blends have become available, all of which pose different challenges for
today’s refiners. Some crudes are both higher in contaminant levels and have a composition that
makes them more difficult to upgrade. Coupled with the demand for increased production of
ultra-clean diesel and gasoline, innovative refinery configuration changes will be needed to
accommodate these new feedstocks. The potential processing schemes under consideration range
from simple hydrotreating for contaminant removal, to hydrocracking and fluid catalytic
cracking for conversion of gas oil to high-quality transportation fuels. It is the integration of
these process technologies, however, that offers the greatest economic potential. This paper
focuses on the processing of heavy and synthetic crude blends using innovative process
integration across several technology platforms to produce clean fuels.

MARKET SITUATION
World oil demand is projected to continue increasing, at a rate of about 1.5% per year, with
increased growth of transportation fuels coupled with a relatively flat heavy oil demand. The
Energy Information Administration recently predicted that the demand for crude oil in the United
States will increase at an average rate of 1.1% through to the year 20302. Most of the 400,000-
barrel annual increase in crude consumption in the United States will be used for transportation
fuels, with gasoline representing about 45% of total petroleum consumption. The demand for
distillate fuels is growing at a faster rate than the demand for gasoline. These changes in demand
AM-06-16
© 2006 UOP LLC. All rights reserved.
Page 1
have been accompanied by well-known ultra-low-sulfur regulations for diesel and gasoline,
forcing a significant capital investment through the end of decade.

The production of light and medium crudes from the Western Canadian sedimentary basin is
dropping by about 20% every five years. Given the increasing demand for oil, and declining
production of conventional crudes in both the United States and Canada, the Canadian oil sands
will become a key source of future crude supply. The volume of oil in place in the various
deposits is estimated at 1.7 trillion barrels, of which 174 billion barrels is recoverable with
existing technology. This places the size of the established reserves second only to Saudi
Arabia. As recovery technologies improve, the size of those recoverable reserves could increase
significantly. The proximity of the source, security of supply, and competitive pricing will drive
the refinery investment needed to accommodate these crudes.

Today, the level of production has reached one million barrels per day, and it is expected to
increase to almost three million barrels per day by 2015. Oil sands crudes are expected to
represent more than 75% of the crude produced in Western Canada. While most of the imports
into the United States are in PADDs II and IV, further pipeline expansions will increase
penetration in PADD II and the Pacific Northwest (Northern PADD V). New pipeline systems
are expected in the future, reaching Texas (PADD III) and a new sea port in British Columbia.
From this new port, marine shipments to both California (Southern PADD V) and the Far East
are expected.

As will be shown later, the composition and contaminant levels of bitumen-derived crudes does
not make them an easy replacement for conventional crudes, especially since most existing
refineries have limited capacity to accept poorer quality feedstocks. These crudes are
fundamentally different, so refiners need to understand them and also be prepared for the
changes needed to process them.

Bitumen-derived Crudes
The term “synthetic crude” has never been strictly defined, but it has come to mean a blend of
naphtha, distillate, and gas oil range materials, with no resid (1050°F+, 565°C+ material).
Canadian synthetic crudes first became available in 1967 when Suncor (then Great Canadian Oil
Sands) started to market a blend produced by hydrotreating the naphtha, distillate, and gas oil
generated in a delayed coking unit. The light, sweet synthetic crude marketed by Suncor today is
called Suncor Oil Sands Blend A (OSA). Syncrude Canada Ltd. started production in 1978,
marketing a fully-hydrotreated blend utilizing fluidized-bed coking technology as the primary
upgrading step. Today, this product is referred to as Syncrude Sweet Blend (SSB). Husky Oil
started up a heavy conventional crude upgrader in 1990 using a combination of ebullated-bed
hydroprocessing and delayed coking technologies. Their sweet synthetic crude is traded as
Husky Sweet Blend (HSB). The Athabasca Oils Sands Project (AOSP) started producing a
sweet synthetic crude in 2003 called Premium Albian Synthetic (PAS) using ebullated-bed
hydroprocessing technology. There are also small volumes of two synthetic crudes produced at
the Consumers’ Co-op refinery called NSA and NSB.

The quality of the kerosene and diesel in these synthetic crude blends has been a major concern
in the past. Bitumen is itself a very aromatic feed, and the choice of primary upgrading
AM-06-16
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technology has an effect on the final distillate quality. Husky’s HSB and AOSP’s PAS crudes
are produced using an ebullated-bed resid hydrocracking technology, and hence have full-range
diesel cetane numbers above 40. Suncor’s OSA crude is produced using delayed coking, giving
the full-range diesel a cetane number very close to, or just below 40. Syncrude’s SSB crude is
produced using the more-severe fluidized-bed coking technology, and the diesel has a cetane
number of about 33. The kerosene cut of SSB has a smoke point of just 13 mm. Starting in
June of 2006, Syncrude will start to produce Syncrude Sweet Premium (SSP), where the
distillate has been further upgraded to give a full-range diesel cetane number of 40, and a
kerosene smoke point of 19 mm.

Table 1 compares the basic composition and quality of Syncrude SSB against Brent crude3. The
SSB crude has lower sulfur and no resid (as shipped), it contains significantly less naphtha-range
material, and more distillate and VGO. Note that due to pipeline contamination, synthetic crudes
like SSB can have some resid component when actually received in the refinery.

Table 1 – Comparison of Syncrude SSB and Brent Crudes

SSB SSB
Brent Produced Received*
Gravity, API 38.6 31.8 33.5
Sulphur, wt% 0.29 0.1 0.2
RVP, psi 8.2 4.6 N/A
Yield, vol%
C4- 2.9 3.4 N/A
C5- 350°F 28.6 15 21
350 - 650°F 29.6 44 40.5
650 - 350°F 29.8 37.6 31.5
1050°F plus 9.1 0 3
Properties
Kerosene smoke pt, mm 25 13 ---
350 - 650°F
Cetane 50 30 ---
Pour -5 -55 ---

While sweet synthetic blends make up the majority of the synthetic crudes on the market, some
sour synthetic blends are also available. Suncor markets a range of sour synthetic blends, each
tailored to meet specific refinery processing capabilities. Suncor OSE crude is a blend of
hydrotreated coker naphtha with non-hydrotreated coker distillate and coker heavy gas oil.
Suncor’s OSV crude is a blend of hydrotreated coker naphtha with straight-run distillate and
straight-run VGO. There are several other sour blends available (OSH, OCC, etc), each with its
own processing characteristics. While these sour crude blends still contain no resid fraction, they
are generally sold to medium and heavy sour crude refineries. Note that AOSP markets a heavy
sour blend called Albian Heavy Synthetic (AHS) which is a blend of their sweet PAS crude with
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unconverted oil from their ebullated-bed resid hydrocracking unit. Since this crude does have a
resid component, it is not really a synthetic crude as it has been defined here, but neither is it a
bitumen blend, like those described below.

Together, the various synthetic crudes make up the majority of the bitumen-derived crudes on
the market today. However, over 400 KB/D of bitumen are produced and shipped to market
without having been upgraded. The bitumen must be diluted with a lighter hydrocarbon stream
to meet the specifications required for shipping in pipelines. DilBits are blends of bitumen and
condensate, typically natural gas condensate. They normally contain 25 – 30 lv% condensate
and 70 – 75 lv% bitumen. The most common streams are Cold Lake Blend (CLB), Bow River
(BRH), and various Lloyd blends (LLB, LLK, WCB). Since the majority of condensate is C5 to
C12 material, and the majority of bitumen is C30+ boiling range material, these crudes have
become known as “dumbbell crudes.” There is a lot of material boiling at each end of the
boiling point curve, but little in the middle.

Natural gas condensate is in short supply in Northern Alberta where the bitumen is produced.
Condensate sells at a significant premium to light sweet crudes for this reason, and some
condensate is actually being shipped by rail back to Alberta from the United States. To address
the shortage of diluent, and the problem with dumbbell crudes, producers have started to market
SynBits, blends of sweet synthetic crude (typically OSA) and bitumen. SynBits have a more
continuous boiling point curve than DilBits, with a significant portion of distillate-range material
in the blend. However, since the synthetic crude diluent has a much lower API gravity than
condensate, more diluent is needed, so SynBits are typically 50 lv% synthetic crude and 50 lv%
bitumen. The most common SynBits on the market today are Christina Lake Blend (CSB) and
MacKay Heavy (MKH), both of which are blends of bitumen produced by Steam Assisted
Gravity Drainage (SAGD) and OSA crudes.

SynDilBits are actually blends of condensate, hydrotreated synthetic crude, and bitumen. They
typically contain about 65 lv% bitumen, with the remaining volume split between the two diluent
streams. The most common of these streams are Wabasca Heavy (WH) and Western Canadian
Select (WCS).

Light sweet synthetic crudes, heavy sour synthetic crudes, DilBits, SynBits, and SynDilBits all
target different refineries. Figure 14 on the next page summarizes how bitumen is upgraded, or
blended with other streams to make various types of crude blends. Light sweet synthetic crudes
are usually sold to light crude refineries, while heavy sour synthetic crudes and DilBits are
normally processed in heavy crude refineries. SynBits and SynDilBits are usually sold to
medium crude refineries, or blended with additional synthetic crude for processing in a light
crude refinery.

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Figure 1 – Disposition of Bitumen-derived Crudes

Refining
Condensate

DilBit Heavy
Crude
Heavy Refineries
SCO SynDilBit
Bitumen SynBit Medium
Production Crude
Refineries
Light
(Mining or In-Situ) SCO SCO/SynBit Blend
Light
Upgrading Crude
Light Refineries
Synthetic
Crude Oil
Source: Purvin and Gertz, 2005
(SCO)

WCS is a somewhat unique SynDilBit in that it has a proprietary formula developed by EnCana,
Talisman, Canadian Natural Resources Limited (CNRL), and Petro-Canada. They wanted to
reduce the large number of heavy crudes being marketed from Western Canada, and achieve
consistency in the heavy crude blends being shipped from Canada. Each batch contains
specified amounts of the following crudes; LLW, LLC, CLB, CSB, MKH, and BR6. As such,
each batch contains condensate, hydrotreated synthetic crude (OSA), heavy conventional crude,
medium conventional crude, Cold Lake bitumen, and Athabasca bitumen. Each batch is blended
to meet the following specifications; API gravity of 19 - 22°, carbon residue of 7 – 9 wt%, sulfur
of 2.8 – 3.2 wt%, and a total acid number (TAN) of 0.7 – 1.0 mg KOH/g.

WCS may well become the new marker heavy crude from Western Canada, and efforts are being
made to have it traded on the New York Commodities Exchange. Production of WCS started in
January of 2005, with shipments currently at 250 KB/D, but expected to increase to more than
500 KB/D by 2008.

Each synthetic crude and bitumen blend has its own unique processing characteristics. The
compositions of three conventional light crudes, two heavy conventional crudes, two bitumen
blends, and one synthetic crude are compared and contrasted in Figure 2. Compared to the
marker West Texas Intermediate (WTI) crude, a typical synthetic crude has no resid, 50% more
VGO, 50% more distillate, and only half the naphtha. DilBits and WCS have about three times
the volume of resid material than WTI, 50% more VGO, but only half the distillate range
material, and half the naphtha.

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Figure 2 – Crude Composition

100%

80%
vol-%

60%

40%

20%

0%
WTI Arab ANS Lloyd Bow DilBit WCS SCO
Light Blend River
Resid Vacuum Gas Oil Distillate Naphtha/LPG

PROCESSING IMPLICATIONS

The processing implications to the refinery will be a function of the type of bitumen-derived
crude imported. First and foremost, refiners used to processing WTI will have a three-fold
increase in the amount of resid coming to the refinery if they replace WTI with WCS. Beyond
the implications to the crude and vacuum unit, extra resid conversion capacity will certainly be
required. Refineries already seem to be anticipating this need. There is about 200 KB/D of
incremental delayed coking capacity in various stages of planning in PADD II, PADD IV, and
northern PADD V1. The addition of coking capacity can bring a new set of issues to these
refiners. Not only must the refinery deal with the coke disposal issue, they must deal with the
cracked products a coker generates. Coker naphtha typically has a high sulfur and nitrogen
content, but is also rich in olefins and diolefins. Coker naphtha cannot simply be added to the
straight-run feed to the naphtha hydrotreater protecting the catalytic reforming unit. The coker
distillate is also high in sulfur and nitrogen, and has a very low cetane number due to its high
aromatic content. The heavy gas oil produced in a coker from bitumen has a particularly high
aromatic content, and therefore makes a poor FCC feedstock. Each of these streams will also
contain differing levels of silica from the antifoam agent used in the coke drums which can
poison the catalyst in downstream hydroprocessing units. Coupled with the high VGO
component of all synthetic crudes and bitumen blends, there is a twofold impact on the FCC unit.
Higher throughputs are required to process the additional VGO, and the feedstock is poor
quality.

Refiners used to processing the VGO from a sweet conventional crude, like Western Canadian
Mixed Sweet (MSW) crude, will see a significant shift in their FCC yield pattern if they start to
process VGO from most bitumen-derived crudes. Table 2 shows commercial data from a
Canadian refinery, and the impact of switching their FCC feed from a VGO from 100% MSW
crude, to a VGO from SSB crude3. Conversion in the FCC unit dropped by more than 20% as
the yields of LCO and decant oil tripled.
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Table 2 – FCC Yield Impact of Processing Synthetic VGO

Western
Canadian Synthetic Crude
Yield, vol% FF
Propylene 11.7 6.2
Butylene 14.0 6.9
Gasoline 61.1 51.3
LCO 9.6 27.6
Decant 2.7 9.3
C3 + Liquid Yield 116 110
Properties
Gasoline
RON 91.9 96.5
MON 80.9 84.3

Most refiners will not see such a drastic change, as they will most likely replace only a portion of
their current crude diet with synthetic crudes or bitumen blends. The impact will be closer to
that shown in Figure 3. This graph shows the FCC yields for the VGOs from three crude blends;
100% Brent crude, 75% Brent plus 25% of a sweet synthetic blend, like Syncrude SSB, and
75% Brent plus 25% of a sour synthetic blend, like Suncor OSE. As the sweet synthetic and the
sour synthetic blends are added to the FCC feed, the yields of LPG and gasoline drop off, while
the yields of light cycle oil and slurry oil increase. This has been one of the historical problems
refiners have experienced while trying to process synthetics. The loss of FCC conversion has
adversely affected the value of the synthetic blends. This adverse impact increases dramatically
as the percentage of synthetic crude in the diet increases. All the VGO-range material in OSE
crude is coker heavy gas oil. This underscores the impact coker gas oil has on FCC
performance, so if a refiner is also adding a new coker to convert the surplus resid in a bitumen
blend like WCS, the overall impact on FCC yields could be even more pronounced.

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Figure 3 – Impact of 25% Synthetic on FCC Yields

80%
70% LPG Gasoline LCO CSO
60%
vol-%

50%
40%
30%
20%
10%
0%
Brent Brent + Sweet Brent + Sour Blend
Blend
Crude Blend

Most aspects of the refinery operation will be affected by the shift to synthetic crudes or bitumen
blends. If a refiner decides to import a sweet synthetic crude like SSB, their reformer feed will
become richer. With more naphthenes in the reformer feed, hydrogen production will increase.
They will, however, have limitations blending distillate fuels. Depending on the other crudes
being processed, the refinery could be limited to just 20% SSB in their crude diet if running for
maximum jet production (smoke point limit), or about 35% if running for maximum diesel
production (cetane number limit). These limits will be relaxed when Syncrude starts producing
SSP in June of 2006. There are also fewer distillate blending constraints with other sweet
synthetics like OSA, HSB, or PAS. The production of ULSD is more difficult than indicated by
the low sulfur level of the distillate cuts of the crude. The sulfur and nitrogen species left in the
kerosene and diesel cuts are the most refractory, difficult-to-treat species that could not be
removed in the upgrader’s relatively high-pressure hydrotreaters. FCC conversion and gasoline
yield will drop significantly when using any of the currently-available synthetic crudes, and the
production of lube base stock may be impossible due to the aromatic nature of the synthetic
VGO. The large percentage of VGO-range material in these crudes may result in the FCC unit
capacity becoming a bottleneck. Since synthetic crudes contain no resid, they can be used to
increase refinery throughput without having to increase resid conversion capacity.

If a refiner chooses to process one of the sour synthetic crudes, all of the benefits and limitations
discussed above for a sweet synthetic crude still apply, but much more severe hydrotreating will
be required at the refinery to produce ULSD or acceptable-quality FCC feed. The sulfur and
nitrogen content of these crudes is very high compared to a conventional light sweet crude.
Additional sulfur plant capacity may be required. The TAN number of these crudes may also
become an issue. Crudes with straight-run VGO components from bitumen, like OSV, have
higher TAN numbers. Metallurgy upgrades may be required to handle these crudes. Sour
synthetic crudes produced from coker products are more aromatic in nature, but the TAN
numbers are much lower.

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DilBits have both a high resid content and a high sulfur content. They are relatively low in
distillate content, and high in VGO content. Significant resid conversion capacity and
hydrotreating capacity would have to be added to a light sweet crude refinery to process a DilBit.
The FCC unit will again be limited by the large volume of the FCC feed. Although better quality
than a coker-derived VGO, the straight-run VGO from bitumen is still very aromatic, and makes
a poor FCC feed. The TAN number of these crudes can be high, depending on the source of the
bitumen. For example, Athabasca bitumens are reported to have higher TAN numbers. The
large volume of condensate in the blend may pose a problem for some refiners. All the light
naphtha may not be able to be blended into the gasoline pool without exceeding RVP
specifications. As already mentioned, some refiners are already shipping the condensate back to
Alberta by rail, both to alleviate any light ends constraint, but also to realize the higher netbacks
for condensate. Depending on the gasoline to diesel (G/D) ratio of the refinery involved, there
may be insufficient material for the distillate pool.

SynBits generally have a lower resid content than DilBits, but more distillate and VGO-range
material. Since the majority of the distillate comes from a sweet synthetic crude, with the
remainder from bitumen, the distillate quality is only marginal. Some hydrotreating will be
required to achieve a 40 cetane number in the full-range diesel. The VGO is still a poor-quality
FCC feed due to the high aromatic content. The TAN number will be lower than that found in a
DilBit since about half the feed has already been severely hydrotreated. The sulfur content of a
SynBit is also much lower than that of a DilBit.

REQUIRED PROCESS MODIFICATIONS

Changes will be needed in a refinery to accommodate the processing of bitumen-derived crude

blends. The addition of resid conversion capacity, of which coking is the most common, to
convert the large volume of resid present in a DilBit, SynBit, or SynDilBit like WCS, is almost a
given. As mentioned previously, there are several new coking projects being considered for that
portion of the refining market currently able to access bitumen-derived crudes via pipelines. The
heavy gas oil produced in the coker, coupled with the large volume of low-quality VGO in these
crudes, overwhelms FCC units designed for VGO from light sweet crudes. These VGO streams
can be upgraded to premium-quality FCC feeds with hydrotreating or hydrocracking. To avoid
any regret capital, one must first have a basic understanding of FCC chemistry, and how that
chemistry is affected by the addition of bitumen-derived feeds to the FCC unit.

The aromatic nature of bitumen-derived VGOs makes them poor FCC feedstocks. The quality of
three bitumen-derived VGOs are compared to the VGO from an Arab Light crude in Table 3.
The bitumen-derived fractions have lower API gravities and lower hydrogen contents, consistent
with higher levels of sulfur, nitrogen, and aromatics. A brief overview of the chemistry of
aromatic conversion in an FCC unit is useful to give more insight into reasons why these
feedstocks are more difficult to process.

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Table 3 – Bitumen-derived VGO Quality

Lt. Arabian Cold Lake Athabasca

Property VGO VGO VGO Coker HGO

Gravity, °API 22.7 16.3 14.2 11.4

Sulfur, wt-% 2.10 3.10 3.77 4.99
Nitrogen, wppm 820 1380 1854 4032
Hydrogen, wt-% 12.47 11.38 10.80 10.43
Carbon Residue, wt-% 0.50 0.67 0.50 0.35
Total Aromatics, vol-% 58 60 69 76
Sim Dist D-2887, °F
IBP / 5wt-% 382/609 597/647 621/693 457/605
10 701 671 719 649
30 777 734 779 719
50 834 795 831 764
70 895 856 878 812
90/95 981/1020 942/980 947/972 878/905
EBP 1100 1054 1032 955

Figure 4 illustrates how a three-ring aromatic compound with several alkyl side chains reacts in
an FCC unit. Methyl and ethyl groups will tend to stay attached to the aromatic compounds.
Alkyl side chains with a carbon number of three or greater will cleave off close to the aromatic
ring. The removed alkyl side chains will initially become olefins, and may crack again into
smaller components depending on the length of the chain. Paraffinic compounds with a carbon
number of five or less tend not to crack. Paraffinic compounds with a carbon number of six
crack slowly, and paraffinic compounds with a carbon number of seven or more will crack fairly
quickly.

FCC cracking will not open the aromatic ring structures. Depending on the number of short
alkyl side chains remaining, the compound in Figure 4 could end up in the FCC light cycle oil
product, but would most likely be produced as part of the heavy cycle oil pool or the decant oil.

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Figure 4 – FCC Conversion of Multi-ring Aromatic Compounds

R1
C
C C2 C2

C FCC Cracking
C
C2 R2 C2
+
R1=C
+
R2=C

If that same three-ring aromatic compound was partially saturated in an FCC feed pretreater
before being fed to the FCC unit, the resulting products would be significantly different (Figure
5). Notice it takes only four moles of hydrogen to saturate two of the three aromatic rings.
Saturated ring structures crack open far more easily in an FCC unit than aromatic ring structures.
While some of the compounds would follow the upper path and partially dehydrogenate back to
a two-ring aromatic compound, the majority of these partially saturated ring structures would
follow the lower pathway. The longer alkyl side chains would first be cleaved off as before. The
saturated rings would then crack open, leaving a single-ring aromatic structure and other
paraffinic and iso-paraffinic molecules.

Figure 5 – FCC Conversion of Partially-saturated Aromatic Compounds

R1 R1
C C g
a c kin C
C C2 + 4H2 C C2
C Cr 2
+
FC
Hydrotreating R1 = C + R2 = C + i-C5
C C
C R2 C R2
C2 C2

FC
CC
rac
ki ng
C2
+
R1 = C + R2 = C + i-C4 + C6=
The single-ring aromatic structures created in this way are almost never benzene. The structures
formed will usually be toluene and xylenes, and thus have high octane number. The iso-paraffins
created will continue to crack to lighter compounds as dictated by the carbon number.

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If that same three-ring aromatic compound was now fully saturated in an FCC feed pretreater
before being fed to the FCC unit, the resultant products would again be significantly different
(Figure 6). It would take seven moles of hydrogen to fully saturate the three-ring structure. As
before, some of the structures would partially dehydrogenate back to a single-ring aromatic
structure, generating several paraffinic molecules. The majority of these saturated ring structures
would crack open completely, generating numerous normal paraffins, iso-paraffins, and olefins.
Again, the paraffins would continue to crack to lighter compounds, depending upon their carbon
number. Normal and iso-paraffinic molecules have much lower octane numbers than aromatic
compounds with the same carbon number. By fully saturating the initial compound in the FCC
feed pretreater, the gasoline yield would be reduced in favor of a slightly higher LPG and gas
yield, and the octane number of the gasoline created would be lower.

Figure 6 – FCC Conversion of Fully-saturated Aromatic Compounds

+ R1 = C
R1 R1 C2
C C i ng +
C C2 + 7H2 C C2 rack
C C R2 = C + i-C5 + C5=
FC
Hydrotreating
C C
C R2 C R2
C2 C2

FC
C Cr
ack R1 = C + R2 = C +
i ng i-C5 + C4= + C5=
+ C3 + C3=

In general, the FCC cracking of saturated and aromatic ring structures can be summarized by
Table 4 below. Cracking of multi-ring aromatics produces high cycle oil yields but low gas,
LPG and gasoline yields. The gasoline produced would consist mainly of paraffins and olefins.
FCC cracking of single-ring aromatic compounds produces low gas, LPG and cycle oil yields,
but a high gasoline yield. The gasoline produced is mostly aromatic, and therefore has a high
octane number. FCC cracking of saturated ring structures produces a low yield of cycle oil, a
reasonably high gasoline yield, but a higher yield of gas and LPG than cracking of single-ring
aromatics. The gasoline produced would again consist mainly of paraffins and olefins, and
would therefore have a lower octane number.

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Table 4 – FCC Cracking of Aromatic Rings (Summary)

High cycle oil yield

Low gasoline yield
Multi-ring Aromatics Gasoline mostly paraffins and olefins
Low gas yield

High gasoline yield

Mono-ring Aromatics Gasoline mostly aromatic (high octane)
Low gas yield

Reasonably high gasoline yield

Naphthenes Gasoline mostly paraffins and olefins
(lower octane)
Higher gas yield

From this, it can be concluded that if the refiner’s objective is maximum gasoline yield from his
FCC unit, the feed pretreater conditions should be set to ensure that all the multi-ring aromatic
compounds are saturated down to single-ring aromatics. Hydrogen addition beyond this point
would only serve to lower gasoline yield and gasoline octane. Higher conversions would, of
course, be achieved, but product value would not be maximized. If a refiner were trying to use
his FCC unit to produce olefins for alky feed or MTBE or other chemical production, obviously
a higher feed hydrogen level would increase the production of olefins.

The curve shown in Figure 7 illustrates the FCC gasoline yield as a function of the feed
hydrogen content. VGOs from more aromatic crudes are near the bottom left-hand end of the
curve, while VGOs from more paraffinic crudes would be located near the top right-hand end of
the curve. The trend shows that gasoline yield increases with increasing feed hydrogen. Gasoline
yield will not continue to rise, however, as the feed hydrogen content is increased. At some
point, the gasoline yield will drop off, as the conversion to gas and LPG continues to increase.
Overall, conversion will continue to increase, but the gasoline yield will drop. This drop off is
not a detriment if the refiner is trying to make olefins for alky feed or petrochemical feedstock
with the FCC unit, but it is detrimental if the FCC unit is designed to achieve the maximum
gasoline yield. This yield response is represented by the dashed portion of the curve, although
the exact inflection point is a function of the feedstock and other variables.

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Figure 7 – FCC Gasoline Yield Controlled by Feed Hydrogen Content

72
Vacuum
70 Gas Oil
68
Gasoline Yield, vol-%
66
64 LC-Finer Hydrocracker
Gas Oil Bottoms
62
60 Coker
58 Gas Oil
56
54
52
50
48
46
11.4 11.6 11.8 12 12.2 12.4 12.6 12.8 13 13.2 13.4 13.6 13.8
Feed Hydrogen Content, Wt-%

Yui et al5 took several intermediate VGO streams derived from Athabasca bitumen, hydrotreated
them at a constant severity, and then processed them in an FCC pilot plant to determine product
yields and qualities. The gasoline yield data from that study is also shown in Figure 7. Note that
the hydrotreated coker gas oil had a hydrogen content of 11.5wt%, and gave a gasoline yield just
above the UOP curve. The gas oil produced in the ebullated-bed resid hydrocracking unit had a
higher hydrogen content after hydrotreating, and generated a gasoline yield above the curve,
while the hydrotreated VGO distilled directly from bitumen produced a gasoline yield well
above the curve. There is a valid reason as to why these bitumen-derived VGO streams all
generated FCC gasoline yields above the standard UOP curve. One of the characteristics of
these VGOs is the very low paraffin content. Some bitumen derived coker VGO samples
analyzed at UOP have actually been shown to contain no paraffins of any kind. All the
molecules were aromatics, naphthenes, or sulfur species which were almost all aromatic
compounds. Without the paraffins to crack to gas and LPG, bitumen-derived VGOs can produce
very high gasoline yields, if the hydrogen content is raised to the appropriate level.

Many people believe that the unconverted oil from a hydrocracker is the best FCC feedstock
available. The Yui study obtained a sample of hydrocracker bottoms from a Canadian refinery
when they were processing 100% synthetic heavy gas oil in their hydrocracker. It is clear that
very high conversions and gasoline yields can be achieved by first hydrocracking the bitumen-
derived heavy gas oil streams. Since the data points fall to the right of the conventional crude
curve, it suggests that hydrocracker bottoms are less efficient gasoline feedstocks than
conventional crudes. The important question becomes, “Is the extra hydroprocessing severity
required to raise the hydrogen content from about 12.3 wt% all the way to 13.7 wt% worth an
extra five or six percent in FCC gasoline yield?” UOP would suggest that hydrogen is most
effectively used to increase the hydrogen content to the 12.3 to 12.5 wt% range. Beyond this
level, more and more naphthenes and paraffins would be created in the FCC feed, decreasing the
gasoline yield and increasing the LPG and gas yield. Since LPG has a high value to many
refiners as alky feed or chemical plant feedstock, the most economic hydrogen content will be
different for each application, and must be carefully determined.

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Table 5 – Required FCC Feed Pre-treat Severity

Source of VGO or HGO 12.0% Hydrogen 12.5% Hydrogen

Vacuum Column 1400 psig 1600 psig
1.6 LHSV 0.9 LHSV
1600 psig 1800 psig
Ebullated-bed Processing
1.2 LHSV 0.8 LHSV
1700 psig 1900 psig
Delayed Coker 1.2 LHSV 0.8 LHSV
1900 psig 2100 psig
Fluidized-bed Coker 0.8 LHSV 0.7 LHSV

Recognizing that the optimum hydrogen content for bitumen-derived heavy gas oil streams may
be between 12.0 and 12.5 wt%, UOP has estimated a rough set of operating conditions necessary
using UOP’s Unionfining™ process technology to increase the hydrogen content of each of the
raw feeds to those levels. As might have been expected, the vacuum heavy gas oil would require
the least severe conditions, where heavy gas oil product from the fluidized-bed coking unit
would require the most severe conditions. These pressure and liquid hourly space velocity
combinations are thought to be close to the economic optimum, but the same results can be
achieved at lower pressure if the space velocity is lowered sufficiently.

Many refiners use the UOP K Factor to estimate the hydrogen content of an FCC feedstock. The
UOP K Factor works very well when trying to distinguish between aromatic and paraffinic
VGO's. However, when the FCC feed is hydrotreated, the UOP K factor-hydrogen content
relationship begins to break down. Thus, if the FCC feed has been severely hydrotreated, the
refiners should focus on API gravity, nitrogen content, and most importantly, hydrogen content
of the FCC feed.

The penalty associated with the quality of the VGO from a bitumen-derived crude changes
rapidly once the FCC feeds have been hydrotreated. Figure 8 summarizes the FCC yield pattern
for the same Brent/synthetic VGO blends shown in Figure 3, but after the feeds had been
hydrotreated to a constant sulfur content. The yield of LPG and gasoline is up in all cases, with
the expected reduction in light cycle oil and slurry oil yields. The magnitude of this yield shift is
not consistent between cases. There is a much greater increase in LPG and gasoline yield with
synthetic VGOs in the FCC feed, especially the sour synthetic blend. When this same analysis
was completed substituting the Brent crude component with the more aromatic Arab Light crude,
the increases in LPG and gasoline yields were even greater, to the point that two of the three
cases had gasoline yields greater than the corresponding Brent crude cases. The most difficult
to treat feeds gained the most advantage by hydrotreating.

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Figure 8 – FCC Yields on Hydrotreated VGO Feeds

80%
80%
70%
70% LPG
LPG Gasoline
Gasoline LCO
LCO CSO
CSO

60%
60%
50%
50%
vol-%

40%
40%
30%
30%
20%
20%
10%
10%
0%
0%
Brent
Brent Brent
Brent +
+ Sweet
Sweet Brent
Brent +
+ Sour
Sour Blend
Blend
Blend
Blend
Crude
Crude Blend
Blend

In 2000, UOP completed an internal study that compared the economics of FCC feed pretreating
with post-treating of FCC products. UOP studied the impact of blending 25% of two different
synthetic crudes into a conventional crude diet. The FCC yield patterns discussed above were
part of that study. One of the conclusions from that study was confirmation that the economics
for pretreating the FCC feed improved as the quality of the raw FCC feed became worse. The
incremental operating costs, capital costs, and net present values for each case are summarized in
Figure 9 below. While all the capital and product prices used for this study are now six years out
of date, the conclusion is still valid; the tougher the feed, the higher the NPV.

Figure 9 – FCC Feed Pre-treat Economics

700
Inc Op Costs, $MM/yr
600
$MM (Year 2000)

ISBL Costs, $MM

500 NPV, $MM

400 Feed V
ghe
r NP
300 u
To r
200 ghe
i
H
100

0
-100
Brent Brent + Brent + Arab Light Arab + Arab +
Sweet Syn. Sour Syn. Sweet Syn. Sour Syn.
Blend Blend Blend Blend

AM-06-16
Page 16
The economics of FCC feed pretreating projects are driven by two factors: better yields from the
FCC unit itself, and incremental conversion. Whenever a VGO-range feed is severely
hydrotreated, about 10 – 15 lv% of the feed is converted to distillate and lighter products. Some
of this conversion is simply a shift in the boiling point curve as sulfur and nitrogen is removed
from the molecules, but some hydrocracking to lighter products does occur. Assuming the
distillate and lighter products are fractionated from the hydrotreated VGO before it is fed to the
FCC unit, hydrotreating the feed effectively debottlenecks the FCC unit. If a refinery has an
FCC unit sized for 50 KB/D, after hydrotreating, there will only be about 43 KB/D of
hydrotreated VGO left to feed to the FCC unit. This allows the refiner to acquire additional
crude if the other processing units have surplus capacity, or at a minimum, it allows the refiner to
purchase surplus VGO to keep the FCC unit full. In either case, it is the incremental conversion
in the FCC feed pretreater that makes the economics so attractive.

UNICRACKING™ PROCESSES
As previously mentioned, all the sweet and sour synthetic crudes, as well as the bitumen blends,
have a high VGO content relative to conventional crude. Hydrotreating alone can turn these
low-quality VGOs into premium FCC feeds and reduce the volume of the feed available, but
most refiners importing bitumen-derived crudes for the first time will still have surplus VGO that
cannot be processed in their FCC unit. This problem will be made worse by the addition of the
heavy gas oil generated in the coker, for those refiners choosing to process a bitumen blend.
Additional VGO conversion capacity must be added. Refiners may consider an FCC expansion
project, or even an entirely new FCC unit, but since they are probably going to construct a high-
pressure hydrotreater to process the FCC feed, it only makes sense to consider converting the
incremental VGO barrels with hydrocracking technology. Hydrocracking has the advantage in
that, unlike an FCC unit, it can be used to produce both gasoline and high-quality distillate fuels.

In order to avoid increased utility cost and unnecessary quality give-away caused by excess
hydrogen consumption, efficient hydrogen consumption is a critical parameter in
hydroprocessing unit design and operation. With growing demand and more stringent
specifications for fuels, it is recognized by industry experts that hydroprocessing technologies
will be key in the future to meeting the refinery conversion capacity and quality needs. Recent
advances in UOP’s hydrocracking technology portfolio, such as the Advanced Partial
Conversion Unicracking (APCU) process, and the other Unicracking configurations discussed
below, were designed for optimal treatment of the distillate and unconverted product fractions,
resulting in more efficient hydrogen utilization.

Once-Through Unicracking Unit

When most refiners think about partial-conversion hydrocracking, they think of a simple once-
through Unicracking process unit. This configuration has the lowest capital cost of the various
Unicracking process configurations. The unit shown in Figure 10 is sized for a fresh feed rate of
60 KB/D with an unconverted oil rate of 30 KB/D, meaning the unit is designed to give a gross
conversion of 50 lv%. If the cracking reactor was loaded with a high-activity naphtha catalyst,
such as UOP’s HC™ 29 catalyst, the yield pattern would be very close to that shown in Figure
10. The unit would produce about 5 KB/D of naphtha, 12 KB/D of kerosene with a relatively
low smoke point, and 17 KB/D of heavy diesel with a reasonable cetane number. Note that if the
AM-06-16
Page 17
kerosene and heavy diesel were combined to form a full-range diesel, the cetane number would
be close to 40.

One key consideration with this simple configuration is the quality of the unconverted oil
produced as FCC feed. Unconverted oil from this unit would have high hydrogen content,
similar to the 13.7 wt% seen in the unconverted oil from the Canadian refinery discussed earlier.
Again, if a refiner is trying to make alky feed or petrochemical feedstock with their FCC unit,
this would make an excellent feed, but if the FCC unit is being run for maximum gasoline
production, the hydrogen content of the unconverted oil may be too high.

Figure 10 – Once-through Unicracking Unit

H2
R-2
Naphtha
61.5 5
R-1
60
Kerosene
12 16 mm Smoke Point
64

Heavy Diesel
17 45 Cetane No

Unconverted
30 Oil

Feed

Separate-Hydrotreat Unicracking Unit

The quality of the distillate fuels produced in a low conversion once-through Unicracking unit
can be poor. Figure 11 is a plot of both kerosene smoke point and diesel cetane index plotted
against conversion in the Unicracking unit. While these plots are based on a VGO from
conventional Arab Light crude, the message is the same for bitumen-derived VGOs. Kerosene
smoke point and diesel cetane index both increase with increasing conversion. If a refiner wants
to generate high-quality distillate fuels from a poor-quality VGO stream, they need to run the
Unicracking unit at a high conversion level.

AM-06-16
Page 18
Figure 11 – Distillate Quality as a Function of Conversion

30 80
Kerosene Smoke Point, mm
25 t
Poin 70

Diesel Cetane Index

ke
Smo
20
60
ndex
15
Cet ane I
50
10

5 40

0 30
0 20 40 60 80 100
Conversion, wt-%
Conventional VGO Feed

Both the distillate quality and the unconverted oil quality can be controlled by using a different
Unicracking unit configuration. In UOP’s separate-hydrotreat configuration shown in Figure 12,
the severity in the R-1 pretreat reactor would be set to achieve the desired hydrogen content in
the unconverted oil, in this example a hydrogen content of 12.5 wt%. The R-1 effluent would be
routed directly to the product fractionator. About 30 KB/D of unconverted oil would be fed to
the cracking reactor, this time running at 80 lv% crack-per-pass. A bed of pretreat catalyst could
be added to the cracking reactor to get the organic nitrogen down to the optimum level for the
cracking catalyst, about 100 wppm or less. The effluent from the cracking reactor would also be
routed directly to the product fractionator.

The fractionator in the separate-hydrotreat configuration would be larger than that in the once-
through configuration, but the high-pressure equipment in the cracking reactor loop would be
50% smaller, now sized for only 30 KB/D. At the high conversion level of 80 lv% crack-per-
pass, the unit would generate higher yields of both naphtha and kerosene, but less heavy diesel.
The quality of both distillate streams would be significantly higher than the once-through unit.
Only about 3 KB/D of the 30 KB/D of unconverted oil produced in the unit would have been
processed to a high hydrogen content in the cracking reactor, thereby reducing overall hydrogen
consumption. The separate-hydrotreat configuration effectively allows independent control of
both the unconverted oil quality and the distillate fuel quality.

AM-06-16
Page 19
Figure 12 – Separate-Hydrotreat Unicracking Unit

H2
R-2
Naphtha
61.5 5
R-1
60
Kerosene
12 16 mm Smoke Point
64

Heavy Diesel
17 45 Cetane No

Unconverted
30 Oil

Feed

Advanced Partial Conversion Unicracking (APCU) Unit

Typical once-through hydrocrackers operating at low to moderate conversion levels may not
produce USLD-quality distillate. The unit pressure is often set by the need to produce low-
aromatic, high-cetane diesel. This higher design pressure results in the production of
unconverted oil with a high hydrogen content, and a higher overall hydrogen consumption. In
the APCU unit flow scheme depicted in Figure 13, the unit is designed to produce ULSD and
partially hydrotreated FCC feedstock as primary products, while operating at a significantly
lower pressure. The refiner achieves two goals with the addition of this unit; increased
production of ULSD, as well as improved quality of the FCC feedstock.

Figure 13 – Advanced Partial Conversion Unicracking Unit

H2
Raw Co-feed
Feed
AMINE

HT PT
Rx Rx
Naphtha
SEP
HC E F.G.
Rx H
S
ULSD

FCC Feed
AM-06-16
Page 20
The benefits of the APCU process compared to a conventional low-conversion hydrocracking
process are reflected in more efficient use of hydrogen due to lower pressure operation, and
staging of reactions in a way that minimizes over-processing and product quality giveaway.
Separation of cracked products from the desulfurized FCC feed in the Enhanced Hot Separator
(EHS), followed by an additional Distillate Unionfining process step, allows the distillate quality
to be controlled independently of FCC feed quality. Over-treating of FCC feed is avoided and
excess hydrogen consumption is minimized. The APCU process can achieve hydrogen savings
of five to ten percent compared to a conventional mild hydrocracking scheme. The integration of
a separate Unionfining reactor in the process enables post-treating of other refinery middle
distillate streams. For refiners processing a bitumen blend following the installation of a coker,
the coker diesel stream must be severely hydrotreated to meet ULSD specs. With the APCU
process, that stream could be processed as a co-feed in the Distillate Unionfining reactor, using
the hydrogen and heat from the Unicracking reactor, resulting in utility cost savings. The level
of hydrotreating of these streams can be independently controlled to add just the right amount of
hydrogen, an added benefit of APCU process.

The typical operating conditions and the resultant FCC feed quality for FCC feed pretreating,
the once-through Unicracking process, the separate-hydrotreat Unicracking process, and the
APCU Unicracking process are summarized in Table 6. Since the quality of the VGO from an
Arab Light crude is so different from that of a WCS VGO, the operating conditions to treat both
feeds are shown for comparison. Note that this comparison represents model predictions for
representative feed properties and that the units are sized for typical run lengths. The table
should be used as a qualitative guide to differentiate between the various processing options.
This is not meant to be a substitute for a customized hydroprocessing solution.

Table 6 – Hydroprocessing Options Summary

FCC Feed FCC Feed Pressure FCC Feed H2 Cons

Feed Unit Type H2, wt-% Nitrogen, ppm ULSD psig LHSV %FF (SCFB)

Arabian Unionfining 12.8 <1000 No Base Base 85 500

Process

Arabian Once-Through 13.5 <50 Yes Base+600 0.3*Base 50 1150

Unicracking
Process

WCS Unionfining 12.4 <1000 No Base+1000 0.3*Base 80 1000

Process

WCS Once-Through 13.2 <50 Yes Base+1000 0.17* Base 50 1900

Unicracking
Process

WCS Separate- 12.5 <500 Yes Base+1000 0.2*Base 50 1700

Hydrotreat
Unicracking
Process

WCS APCU Process 12.9 <50 Yes Base+800 0.17*Base+ 50 1800

+
Plus additional volume for diesel co-feed
AM-06-16
Page 21
Many refiners believe they can treat a portion of their VGO in a Unicracking unit to generate
some high-quality VGO, and then blend it with lower quality VGO, like coker gas oil, to get an
FCC feed with hydrogen content near the 12.5 wt% UOP believes to be close to the optimum
level. While this may be true mathematically, the FCC unit in actuality is concerned only with
the types of molecules being fed to it, and not the average hydrogen content. The three- and
four-ring aromatics in the coker gas oil will still end up in the cycle oil or the slurry oil, and the
naphthenes and paraffins in the hydrocracker bottoms will still be mostly converted to gas and
LPG. Some hydrogen donor reactions will occur in the FCC unit to transfer hydrogen from the
hydrogen-rich stream to the hydrogen-deficient stream, but the overall yields will still be
significantly worse compared to those obtained if the entire feed stream has been hydrotreated to
a hydrogen content of 12.5 wt%. Some refiners have reported significant synergies when
processing Unicracking unit unconverted oil with lower-quality VGOs in their FCC unit.

EXPERIENCE AND PILOT PLANT STUDIES

UOP has been very successful in licensing technology for synthetic crude projects in Canada, the
United States, and Venezuela. Our understanding of the processing issues is based on extensive
pilot plant and commercial experience. UOP has completed both hydrotreating and
hydrocracking pilot plant studies on numerous distillate and VGO-range feeds derived from
Athabasca bitumen, Cold Lake bitumen, Lloydminster heavy conventional crudes, and Orinoco
Belt heavy crudes.

UOP has licensed 22 of the 32 hydroprocessing units currently in operation, design, or

construction for the upgraders in Western Canada. UOP’s experience with hydrotreating,
hydrocracking, FCC, coking, and reforming gives us the unique ability to look at all the units
affected by import of any of these bitumen-derived crudes.

CONCLUSIONS
The supply of Canadian light and medium sweet crudes is declining. There is a vast oil sands
resource in Canada which up until recently, has not been exploited to any great degree.
Synthetic crudes and bitumen blends will become the dominant crudes in PADD II, PADD IV,
and Northern PADD V in the next ten to fifteen years. Refiners planning to process one of the
bitumen blends will have to install additional resid conversion capacity. Refiners planning to
process any bitumen-based crude will have to install FCC feed pretreating capacity to maintain
acceptable FCC yields. Many refiners will have to install hydrocracking capacity to convert the
large volume of VGO-range barrels present in these crudes, and to meet the increasing demand
for distillate fuels.

The hydrogen content of an FCC feedstock is the critical parameter to control when trying to
optimize the performance of all the VGO processing units. UOP has developed novel
Unicracking process flow schemes that allow the refiner to control the hydrogen content of the
FCC feed independently from conversion level or distillate quality.

AM-06-16
Page 22
REFERENCES
1. Purvin & Gertz Inc., “Global Markets for Canadian Oil Sands Crudes”, December, 2005
2. Energy Information Administration, “The AEO2006 Early Release”, www.eia.doe.gov,
January 2006
3. Halford, T.L., McIntosh, A.P., and Rassmussen, D., “A Canadian Refiner’s Perspective
of Synthetic Crudes”, Proceedings of the 1997 NCUT Conference, Directions in Refining
and Marketing of Synthetic Crude Oil (SCO) and Heavy Oil, September, 1997
4. Purvin & Gertz Inc. “Fuel Options for Oil Sands Development”, September, 2004
5. Yui, S., Matsumoto, N., and Sasaki, Y., “Athabasca oil sands produce quality FCC
feeds”, Oil & Gas Journal, January 19, 1998 Journal
6. www.crudemonitor.ca – a website sponsored by the Canadian Association of Petroleum
Producers

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