As crude oil and product prices continue to climb, there are
economic incentives for refineries to increase the total
distillate yield with increased selectivity toward diesel fuel.
The debate continues over the pros and cons of simply adding a
new delayed coker vs. a residue hydrocracker upstream of an
existing delayed coker to improve overall liquid yield. The
following commercial examples will explore both sides on how to
find more distillates from every barrel of oil processed.
The US has the greatest concentration of delayed cokers in
the world. Of the 130 refineries processing 17.8 million bpd
(MMbpd) of crude oil, 60 of these refineries use delayed coking
(DC) to destroy vacuum residue (VR) and to increase the yield
of distillates for further processing into transportation
The first delayed coker came online in 1929 at the Standard
Oil of Indiana refinery located in Whiting, Indiana. At that time, crude oil was
selling for $1.27/bbl in current dollars. Since then, the refining industry has gone through
various economic cycles. The most recent cycle started in 2000
with a consistent rise in the price of crude oil, which is
about $100/bbl for WTI. In addition, two major shifts have
occurred in the energy market; natural gas prices started
declining in 2008 due to new discoveries, and the
gasoline-to-diesel margin reversed in 2005 with diesel priced
higher than gasoline. As a result of these changes, a study was
conducted to explore how these shifts would influence a
refiners decision on adding more crude capacity via a
US MARKET FOR DELAYED COKING
In 2010, the US had 60 delayed cokers as compared to 11 in
Europe, 4 in the Middle East and 27
in Asia-Pacific. In the US market, delayed coking was the
preferred choice for destroying the VR from medium and heavy
crude oil. Of the 60 delayed coking units in the US, 55% in
terms of capacity are located in PADD 3 (US Gulf Coast) and 13%
are located in the PADD 2 market (US Midwest). The vast
majority of these delayed coking units (DCUs) were installed
when crude oil was below $20/bbl. In the last 10 years, Brent
and WTI prices have risen at an unprecedented rate.
Historically, refineries have added incremental DC capacity
as part of refinery expansions because it was considered to be
low-investment, well-known and economically attractive option.
But, with the new changes in market prices and the increase in
residue hydrocracking worldwide, is DC still the best option
for a US-based refinery?
Case history 1
The following case history investigates an existing
100,000-bpsd refinery processing 100% Arabian Heavy crude.
Expansion studies were conducted using both Arabian Heavy crude
and Athabasca bitumen, with properties as listed in
Fig. 2 is the front-end section of a
typical refinery configuration; it uses a delayed coker to
process the entire VR. Straight-run (SR) and delayed coker
distillates are processed into naphtha, diesel and fluid
catalytic cracking (FCC) feed pretreat hydrotreaters.
Steam-methane reforming (SMR) is used to generate hydrogen.
Fig. 1. Delayed coker at
upgrader at Lloydmininster, Saskatchewan,
Fig. 2. Base case of
Two expansion configurations were investigated for this case
study. The first case (Case 1) adds an additional 100,000 bpsd
of Arabian Heavy crude and duplicates the existing 27,200-bpsd
delayed coker. The expansion brings the total crude throughput
to 200,000 bpsd. The battery limits are shown in Fig.
2; it includes the associated offsite and utilities.
It does not include the FCC unit or the post-FCC
The second case (Case 2) adds a 54,400-bpsd residue
hydrocracker upstream of the existing delayed coker, which
remains unchanged, as shown in Fig. 3. In this
configuration, the residual hydrocracking unit is a
single-train plant with a single reactor operating at 60%
conversion of 975°F+ residue to distillates. A second
variation of Case 2 (Case 2A) was also investigated with a
variation in which the conversion level is increased to 70% and
the crude throughput is increased to 300,000 bpsd to fill the
Fig. 3. Case 2 of the
100,000-bpsd refinery with
residue hydrocracker and expanded DCU.
To handle the increased feedrate and reactor severity for
Case 2A, two reactors, in series with inter-stage separation
was required for this single-train plant. With the higher
conversion level in the residue hydrocracker, the total Arabian
Heavy crude capacity was increased to 300,000 bpsd, and the
existing DCU is still capable of processing the entire
The final case (Case 3) examines the effect of switching
from Arabian Heavy Crude to Athabasca bitumen (DilBit). This is
a variation of Case 2 (see Fig. 3) with the
addition of a residue hydrocracker ahead of the existing DCU.
Due to the high VR content in the crude, the crude rate to the
refinery is only increased from 100,000 bpsd to 150,000 bpsd.
In all cases, the product streamsnaphtha, diesel and
vacuum gasoil (VGO)are treated to the same level of
For this updated study, pricing data from the US Energy
Information Agency (EIA) for the US as a whole and also for the
PADD 2 (Midwest) and PADD 3 (Gulf Coast) were examined. The US
prices for Brent, WTI and industrial natural gas for the last
10 years are shown in Fig. 4. Brent and WTI
have tracked fairly close to each other, except for the last
couple of years. The prices for DilBit (Athabasca bitumen) can
be calculated from Western Canadian Select (WCS) synthetic
crude, which is traded in Chicago, Illinois. There is a weak
correlation between Brent and WCS prices except when the
natural gas condensate (diluent) is removed from the WCS. To
calculate the actual price of the bitumen, the cost of natural
gas condensate is removed from the DilBit resulting in an
average net price for the Athabasca bitumen at $68/bbl when
Brent crude is valued at $100/bbl. This bitumen price is the
same price as Hardisty heavy bitumen (12°API) at
$68.35/bbl, quoted in January 2012.
Fig. 4. Crude oil and
natural gas prices: June
2000 to June 2012.
Brent crude is used as benchmark crude for this study to
determine gasoline and diesel margins based on historical
trends. Natural gas prices increased from 2002 to 2005 due to a
large demand and supply shortage, as shown in Fig.
4. However, in 2006, the production of additional
natural gas entered the market with tight shale gas formations;
this started a downward trend in natural gas prices. At
present, the average US industrial natural gas price is in the
range of $4 to $5 per thousand standard cubic feet (scf), or
roughly $30/bbl of oil equivalent (boe) basis.
The gasoline-to-Brent price spread (gasoline price minus
Brent crude price) is shown in Fig. 5. It
reflects a general increase in gasoline margins from 2000 to
2007 and then a steady decrease thereafter. Importation of
gasoline from Europe and the increase in ethanol blending into the US
gasoline pool are decreasing domestic demand for this fuel. For
the last three years, PADD 2 prices have been higher than the
average US prices while PADD 3 prices have been lower than the
national average. The diesel-to-gasoline margin over the same
period is shown in Fig. 6. For this study, a
price spread of $9/bbl is assumed for gasoline to Brent crude,
which equates to $109/bbl for gasoline when Brent crude is
valued at $100/bbl for the average US market. Slightly higher
prices could be used for projects in PADD 2, based on these
Fig. 5. Gasoline-to-Brent
price for US,
PADDs 2 and 3: 1992 to 2012.
Fig. 6. Diesel-to-gasoline
prices: 1992 to
The price of diesel fuel overcame the price of gasoline in
2005, and it has remained higher than gasoline for the last six
years. Consequently, there is more interest from refiners to
increase diesel production. This would imply an increase in
mild- and full-conversion hydrocracking in the future. Based on
the mentioned trends, an economic basis, as shown in
A summary of the four expansion cases investigated is
Case 1: Add 100,000 bpsd of Arabian Heavy
crude to the existing refinery using DC as the residue
Case 2: Add 100,000 bpsd of Arabian Heavy
crude and install a residue hydrocracker, operating at 60% VR
conversion ahead of the existing delayed coker.
Case 2A: Same as Case 2, with the residue
hydrocracker operating at 70% VR conversion and crude
throughput increasing to 200,000 bpsd.
Case 3: Add 50,000 bpsd to the existing
refinery and switch from Arabian Heavy crude to Athabasca
bitumen. In this case, a residue hydrocracker was installed
upstream of the existing DCU.
In all of the cases evaluated, the SR and cracked products
were hydrotreated to meet the product specifications, as shown
in Table 3.
Expansion Case 1
The existing refinery crude capacity was doubled to 200,000
bpsd with Arabian Heavy crude. The total VR feedrate to the
delayed coker is 54,400 bpsd. The new coker is a duplicate of
the existing unit. The C5+ product yield
from the delayed coker is 66 vol%. This product is then blended
with the SR distillates and hydrotreated to meet the product
specifications, as listed in Table 3. The
overall liquid yield was 180,500 bpsd or 90.3 vol% on crude
throughput, which includes liquefued petroleum gas (LPG),
naphtha, diesel and VGO. The VGO is assumed to be routed to an
FCC unit, which has a post-hydrotreater and can meet Tier 3
gasoline specifications. Table 4 summarizes
the breakdown of the product yields.
Expansion Case 2
As in Case 1, the overall refinery throughput is doubled to
200,000 bpsd and all of the VR (54,400 bpsd) is routed to a
single-train, single-reactor residue hydrocracking unit
operating at 60% VR conversion. The unconverted residue (21,922
bpsd) is sent to the existing DCU with a nameplate capacity of
27,200 bpsd. The overall yields from the residue hydrocracker,
the downstream delayed coker and hydrotreaters are listed in
Table 4. All of the SR, residue hydrocracker
and coker distillates are hydrotreated to meet the product
quality specifications, as shown in Table 3.
The overall liquid yield was 192,600 bpsd or 96.3 vol % on
crude throughput including LPG, naphtha, diesel and VGO, which
is routed to an FCC unit.
This case is very similar to the commercial residue
hydrocracker/DCU at Husky Energys Lloydminster Upgrader
in Saskatchewan, Canada (Fig. 7). The feed to
this residue hydrocracker is about 34,000 bpsd of a blend Cold
Lake/Lloydminster heavy residue, and it operates around 60%
conversion. The entire unconverted residue from the residue
hydrocracking unit is routed to a DCU to produce fuel-grade
coke for export.
Fig. 7. Husky Energys
unit in Lloydmininster,
In this case, the residue hydrocracking unit conversion
level is increased from 60% to 70%. The number of reactors is
increased to two in series with inter-stage separation, but
they still operate in a single train. The larger reactor volume
is required due to the greater feedrate and reactor severity.
With the higher conversion level, the refinery throughput can be increased
to 300,000 bpsd, which results in a feedrate of 81,655 bpsd to
the residue hydrocracking unit; the unconverted bottoms (24,503
bpsd) are routed to the existing DCU. The yields for this case
are shown in Table 4 for the residue
hydrocracker and DCU. As before, all of the distillate SR and
residue hydrocracked/delayed coker products are hydrotreated.
The overall liquid yield is 292,300 bpsd or 97.4 vol% on crude
Expansion Case 3
In this case, the crude type is switched from Arabian Heavy
to a Canadian DilBit based on Athabasca bitumen. The feedrate
to the refinery is expanded to only 150,000 bpsd of Athabasca
bitumen (excluding the diluent, which is recovered and returned
to Canada). The total feed-rate to the diluent recovery unit is
216,900 bpsd, and it contains about 31 vol% of diluent. The
relatively small increase in throughput is due to the high
content of VR in the feed (58.6 vol% vs. 31.9 vol% for Arabian
Heavy). The feedrate to the residue hydrocracking unit is
83,754 bpsd, and the feedrate to the delayed coker is 27,221
bpsd. In this case, the residue hydrocracking unit is a single
train with two reactors in series with interstage separation
and operates at 68% conversion.
A summary of the cases processing Arabian Heavy crude is
shown in Table 4. The most severe design
conditions were associated with the cases processing the
greatest percentage of cracked stocks and the highly aromatic
bitumen feedstock. Catalyst cycle lengths
were set at 30 months. The product naphtha is routed to a
catalytic reforming or isomerization unit, diesel to the
ultra-low-sulfur diesel (ULSD) pool and VGO to the FCC/post
treater for meeting Tier 3 gasoline specifications.
In residue hydrocracking, many of the coke precursors are
hydrogenated, which results in higher liquid yield and reduced
coke production. In addition, hydrogen consumption in the
liquid product increases the API gravity, which, in turn, leads
to greater volume swell and increased yield of transportation
As expected, the total liquid yield is a function of the
residue conversion level and the amount of hydrogen consumed in
the liquid product, as shown in Table 4. Case
2 shows a 6 vol% increase in liquid yield from Case 1, which is
about 4.2 MM bbl/yr of additional product (LPG, naphtha, diesel
and VGO). By increasing the residue hydrocracker conversion
from 60 vol% to 70 vol%, the total yield increases by 6.6 vol%
over Case 1, which adds an additional production of 4.6 million
bbl/yr of liquid product. The additional production translates
into additional net revenue (product revenue less feedstock
cost and operating cost), as shown in Fig. 8.
The Case 1 expansion adds an additional $77 million/yr, while
Cases 2 and 2A add more than $500 million net revenue/yr. In
contrast with the higher liquid yield, the coke production is
reduced by more than 50%. Coke produced in Cases 1 and 2 are
3,114 metric tpd and 1,431 metric tpd, respectively, indicating
that 54 wt% of the coke precursors were converted in the
residue hydrocracker. When the residue hydrocracker conversion
is raised to 70%, the conversion of coke precursors is
increased to 63 wt%, reducing the amount of coke even further.
Accordingly, Case 2A can process more feed without major
modifications to the existing DCU.
Fig. 8. Product yield of
conversion unit vs.
Selectivity to diesel fuel
Ebullated-bed residue hydrocrackers are more selective to
middle-distillate production than other conversion
technologies. With the margin between diesel and gasoline
expecting to increase, the selectivity becomes more important
to the refiner desiring to maximize the economic returns on projects. One measure of this
selectivity is the ratio of diesel-to-gasoline production. As
shown in Fig. 9, the selectivity of the
conversion unit increases from the DC scheme (Case 1) of 1.5
bbl of diesel to 1 bbl of gasoline production to the residue
hydrocracker/DCU (Case 2at 60% conversion) and reaches
the highest value 2.2 bbl of diesel to 1 bbl of gasoline for
the residue hydrocracker/DCU (Case 2Aat 70% conversion).
For a 200,000-bpsd refinery, the diesel production would
increase from 64,400 bpsd (Case 1) to 70,500 bpsd (Case 2A).
The incremental increase of 6,000 bpsd translates into an
annual revenue increase of $234 million for the refinery.
Fig. 9. Diesel/gasoline
selectivity for Cases 1,
2 and 2A.
Hydrogen consumption and volume swell
As shown in Table 4, the total hydrogen
consumption for the expansion increases by 78% from Case 1 to
Case 2A. This results in a total volume swell increase of 6.6
vol% on crude, which equates to an additional product of 13,200
bpsd for a 200,000-bpsd refinery. As mentioned previously, the
base price of industrial natural gas used for this study is
$5/thousand scf, which is about $30/bbl (boe basis). With
gasoline and diesel selling for $109/bbl and $114/bbl, hydrogen
consumption provides the refinery with an impressive uplift of
$79/bbl to gasoline (i.e., $30/bbl H2 boe to
$109/bbl for gasoline) and $84/bbl uplift for diesel
Alternate case when processing Athabasca bitumen
The major results of this case are shown in Table
4. Processing Athabasca bitumen or other heavy
Canadian crudes will provide economic advantages that include
upgrading a cheaper feedstock with low-cost hydrogen to
high °API transportation fuels.
Relative to Case 2A, Case 3 provides the highest total
liquid yield on crude of 101.5% of total liquid product vs.
96.9% for Case 2A. This is due to the lower API gravity of the
crude and upgrading to about the same API gravity of the
products. This case also represents the highest production of
diesel and VGO per barrel of crude for any of the cases
For all of the cases, diesel production could increase
further by adding a VGO hydrocracker during the expansion as
compared to adding additional cat feed hydrotreating (CFHT)
capacity upstream of the FCC unit. This would also improve the
overall refinery diesel/gasoline ratio.
For the economic analysis, the same basis from Table
2 will be used. The investment cost for the expansion
cases was only for new units and associated offsites and
utilities whereas, the revenues and operating expenses were for
the entire refinery.
Offsites and utilities were taken as a percentage of the
total installed cost for the process units. Fig.
10 shows the investment cost breakdown for each of the
investigated cases. The investment cost per barrel of crude for
the new units varied from $14,800/bpsd to $22,400/bpsd with the
delayed coker expansion at the lowest overall investment. For
the expansion cases, the investment cost included new crude and
vacuum units; conversion unit (delayed coker or residue
hydrocracking unit); naphtha, diesel and VGO hydrotreaters, SMR
hydrogen plant, sulfur plant, gas recovery section, amine regeneration;
sour-water stripping; and corresponding offsites plus
Fig. 10. Investment costs
for various expansion
plans: Cases 1, 2, 2A and 3.
Operating cost for ISBL
The total operating cost, including fixed and variable
operating costs, varied from $3.15/bbl of crude in Case 1 to
$4.42/bbl for Case 2. Case 3, processing Athabasca bitumen, was
the highest with a cost of $7.98/bbl. The top two operating
costs for the residue hydrocracking unit/DCU cases (Cases 2, 2A
and 3) were natural gas plus catalyst and chemicals vs. natural
gas and electricity for the delayed coker case (Case 1).
Rate of return
The total net annual revenues (product revenue less crude
cost and total operating cost) varied from $169 million for the
delayed coker expansion (Case 1) to $932 million for the
residue hydrocracking/DCU expansion (Case 2A) based on an
Arabian Heavy crude price of $92.48/bbl. The product prices
were $109/bbl for gasoline and $114/bbl for diesel.
As shown in Fig. 11, the addition of a
residue hydrocracker upstream of a delayed coker is more
profitable when Brent crude price exceeds $55/bbl. As light oil
prices continue to climb, the IRR for the delayed coker
expansion case falls to zero when Brent crude reaches $115/bbl.
This is due to the low conversion (i.e., low product liquid
yield) and high crude costs. This analysis assumes a constant
$/bbl discount to Arabian Heavy crude and a constant $/bbl
differential between the price of gasoline and diesel to the
price of Brent crude. History tells us that variations will
occur in both light- and heavy-crude price differentials as
well as price fluctuations in the finished product prices of
gasoline and diesel. For this reason, several sensitivity
studies were conducted.
Fig. 11. IRR vs. Brent
During a sensitivity study, a number of questions were asked
including, What happens if the diesel-to-gasoline spread
continues to widen? In all cases, the IRR climbs sharply
by 6 to 7 percentage points for every $5/bbl the margin of
diesel/gasoline increases. In the US Energy Information website
forecast, the margins are expected to keep climbing for the
Whats the impact in processing Athabasca bitumen from
Canada relative to Arabian Heavy? The IRR doubles from 24% in
Case 2 to over 50% in Case 3. This is mainly due to the
attractive price of Canadian bitumen ($68.85/bbl) vs. the price
for Arabian Heavy ($92.48/bbl). The differential of $23.63/bbl
for feedstock cost provides a
significant incentive for all cases processing Athabasca
During a review of product prices in the US market, it was
noted that higher margins for diesel fuel in PADD 2 (Midwest
market) were $2/bbl to $3/bbl. The price variations in the
diesel/gasoline spread varied between $5/bbl to +$17/bbl with a
general increase occurring over the past five years. As shown
in Fig. 12, an increase in the price of ULSD
fuel vs. gasoline provides a tremendous uplift in the IRR for
Fig. 12. IRR vs.
A project located in the Midwest would see the IRR increased
by 4 to 6 percentage points, depending upon which expansion case is selected. The same
general trend is evident when the gasoline-to-Brent crude price
Residue hydrocracking based on ebullated-bed technology is a mature technology,
with 17 operating plants processing more than 650,000 bpsd of
VR in North America, Europe, Middle East and
Asia-Pacific. The reliability of the advanced
ebullated-bed technology has improved over the last 44 years
since the startup of the first plant for KNPCs Shuaiba Refinery in Kuwait.1 Over
the past 10 years of operation, the average availability of six
commercial advanced ebullated-bed units was 96% (Fig.
13).1 This high level of reliability is the
direct result of nearly 200-unit years of operating experience,
automation of operations, pro-active reliability teams, improvements in
the understanding of the chemistry of asphaltene conversion and
stability through R&D, and ongoing improvements in critical
equipment, components and process instrumentation. The plot
shown in Fig. 13 reflects unit availability
for six operating commercial advanced ebullated-bed
units.1 Availability is defined as the actual
onstream time less planned turnarounds (typically occur once
every three to six years) and outages due to external factors
(i.e., hurricanes on the Gulf
Fig. 13. Onstream times for
advanced ebullated-bed hydrocrackers in
Jim Colyar, a senior technology
consultant, performed the revised internal study for which this
article is based. The authors wish to acknowledge his work and
contribution to heavy-oil upgrading.
1 The process is
Axens ebullated-bed technology.
Duddy, J., L. Wisdom, S. Kressmann,
and T. Gauthier, Understanding and Optimization of
Residue Conversion in H-Oil, Oct. 20, 2004.
Ellis, P. J. and C. A. Paul, Delayed coking, AIChE
1998 Spring National Meeting, New Orleans, March 812,
Largeteau, D., J. Ross, M. Laborde and L. Wisdom, The
Challenges & Opportunities of 10 wppm Sulfur
Gasoline, 2011 NPRA Annual Meeting, San Antonio, March
McQuitty, B., Status of the Bi-Provincial Upgrader: H-Oil
Operation and Performance, IFP Seminar in Lyon, France
US Energy Information Agency, 2012 Annual U.S. Crude Oil
First Purchase Price.
Wisdom, L., E. Peer. and P. Bonnifay, Cleaner fuels shift
refineries to increased resid hydroprocessing, Parts 1
and 2, Oil and Gas Journal, Feb. 9, 1998.
Wisdom is a senior executive at Axens in charge
of marketing the heavy-ends technologies in North
America. The current portfolio of technologies includes
the hydrotreating and hydrocracking of gasoil and
residues, slurry-phase hydrocracking, solvent
deasphalting and visbreaking. During his 30 year career,
he has co-authored more than 30 papers on heavy-oil
upgrading and holds two patents. Prior to joining Axens,
he worked for Hydrocarbon Research Inc. (HRI) and
FMC Corp. Mr. Wisdom graduated from the University of
Kansas with a BS degree in chemical engineering and a MBA
in marketing and finance.
is the director of heavy oil and coal technology for
Axens North America Inc. in Princeton, New Jersey. He is
responsible for Axens ebullated-bed technologies
for upgrading of heavy oil and coal. These technologies
include H-Oil, H-Coal and Coal/oil co-processing. Mr.
Duddy has been with Axens for 32 years and holds a BS
degree in chemical engineering from Drexel
Frédéric Morel is an
expert director adviser for Axens marketing, technology and tech services
department. He was formerly the manager of Axens
hydroprocessing and conversion technical services group
and the product line manager of VGO, resid and coal
conversion. Mr. Morel has over 30 years of experience
in oil refining, having worked
previously with IFPs Lyon Development Center as a
research engineer, a project leader of
distillates and residues hydroprocessing, and the
manager of the development department. Mr. Morel holds
a degree in chemical engineering from Ecole
Supérieure de Chimie Industrielle de Lyon and a
graduate degree from Institut dAdministration des