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 fuels.
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 refinery expansion.
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 Table 1.
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 Husky Energys
upgrader at Lloydmininster, Saskatchewan,
Fig. 2. Base case of 100,000-bpsd existing
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 hydrotreater.
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 existing DCU.
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 unconverted VR.
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 product quality.
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 historical trends.
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 Table 2,
A summary of the four expansion cases investigated is described here:
Case 1: Add 100,000 bpsd of Arabian Heavy crude to the existing refinery using DC as the residue conversion unit.
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 residue hydrocracking
unit in Lloydmininster, Saskatchewan,
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 throughput.
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 fuels.
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 production.
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 examined.
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 utilities.
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 price.
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 short term.
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 bitumen.
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 the project.
Fig. 12. IRR vs. diesel-to-gasoline spread.
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 is increased.
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 Coast). HP
Fig. 13. Onstream times for commercial
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, 1998.
Largeteau, D., J. Ross, M. Laborde and L. Wisdom, The Challenges & Opportunities of 10 wppm Sulfur Gasoline, 2011 NPRA Annual Meeting, San Antonio, March 2011.
McQuitty, B., Status of the Bi-Provincial Upgrader: H-Oil Operation and Performance, IFP Seminar in Lyon, France September 1997.
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.
||Larry 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.|
||John Duddy 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 University.|
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 Entreprises.