US 20030098258 A1
A method is provided for increasing the net present value of fuel production while reducing the sulfur content of one or more fuel products produced in a refinery. The method includes converting at least one existing reformer into a hydrodesulfurization system; and replacing the at least one hydrocarbon reformer with a continuous catalytic reformer.
1. A method for increasing the net present value of fuel production while reducing the sulfur content of one or more fuel products produced in a refinery having a first production capacity, comprising:
converting at least one hydrocarbon reformer into a desulfurization system; and
replacing the at least one hydrocarbon reformer with a modem reformer, to form a second production capacity;
wherein the net present value of the second production is greater than the net present value of the first production capacity.
2. The method of
3. The method of
4. The method of
5. The method of
6. The method of
7. The method of
8. A method for reducing the sulfur content of diesel fuel and gasoline in a fuel production process to comply with government standards, comprising the steps of:
converting at least two catalytic reformer units into hydrodesulfurization units;
adding a continuous catalytic reformer to the refinery infrastructure; and
wherein the net present value of fuel produced by the fuel production process using the CCR and the desulfurization systems, is greater than the net present value of the fuel produced by the fuel production process using at least two catalytic reformers.
9. The method of
10. The method of
11. The method of
12. The method of
13. The method of
14. The method of
15. The method of
16. A method for producing low sulfur gasoline and ultra low sulfur diesel in a refinery process having a gasoline production stream, a diesel fuel production stream, or both, with said streams having as part of their refining production streams at least one reforming step, comprising:
converting the at least one reforming step of the gasoline production stream into one or more hydrodesulfurization steps, and
replacing the at least one reforming step in the process with a higher efficiency catalytic reforming step, to produce an increase in value of the production stream or streams which offsets at least in part an initial cost for the step of converting the at least one reforming step of the gasoline production stream into one or more hydrodesulfurization systems.
 The present invention relates to the field of petroleum refining. In particular, the invention relates a method of producing low sulfur gasoline (LSG) and ultra low sulfur diesel (ULSD) which maximizes the net present value of the gasoline produced by a refinery.
 The process of refining petroleum into gasoline, diesel fuel and other products is widely practiced throughout the world. In general, this process involves distilling of crude oil into a variety of hydrocarbon fractions, reforming or purifying at least some of the fractions to improve their performance characteristics for their intended use, and, optionally, blending the fractions into marketable fuel products.
 Sulfur containing compounds are significant impurities in both the crude oil and the distillation fractions, and their removal has become particularly important in recent years. New federal environmental regulations have imposed strict standards for lowering the sulfur content of both gasoline and diesel sold and used in the United States. Similar standards are already in place in parts of Europe and in California. In short, to meet these standards, the sulfur content of gasoline must be reduced to 30 ppm (parts per million), while the sulfur content of diesel fuel must be reduced to 15 ppm under the coming regulatory scheme.
 Even without these regulatory requirements, the removal of sulfur impurities is highly desirable for a number of reasons. Such impurities can degrade catalysts and corrode machinery in later reforming/blending steps. Moreover, they may also produce a noxious smell when used in combustion engines, as well as contribute to environmental problems such as acid rain.
 There are a number of technologies that can be used to remove these impurities from the hydrocarbon fractions during the refining process. These technologies usually involve “hydrogenation”, the catalytic addition of hydrogen molecules to the sulfur impurities to form compounds with lower boiling points (e.g., H2S). These compounds can then be removed from the hydrocarbon stream.
 While these technologies are widely known, upgrading a refinery to include units employing them can be extremely expensive, requiring all new reactors, catalyst beds, piping, heaters, towers and other equipment and infrastructure. Also, the construction of new reactors, beds, and the like may represent an inordinate cost to some refinery operators, especially in light of the limited profit margins available in the commercial distribution and sale of gasoline and diesel fuel. As a result, many refineries are economically constrained with respect to their current infrastructures, and capital expenses and costs to implement desulfurization technologies are eventually borne by either the retailers, consumers, or the refineries' shareholders.
 Thus, there is a need in the art for a method of implementing sulfur reducing technologies to produce LSG and ULSD while increasing the net present value of fuel production in current petroleum refineries.
 The present invention relates a method of producing LSG and ULSD while increasing the net present value of fuel production by selectively converting existing reformers within the refinery into hydrodesulfurization units, and replacing the converted reformers units with modern units. This results in an increase in the amount of gasoline produced by the refinery, an increase in the amount of high octane rated content produced by the refinery from the hydrocarbon fractions or streams derived from the crude petroleum, or both. Thus, the present invention allows refinery operators and owners to produce LSG and, optionally, ULSD in a cost effective manner.
 The method involves converting a hydrocarbon reformer into a desulfurization system, and replacing the old reformers with state of the art reforming technology. The net present value of the refinery's production capacity increases relative to its first production capacity (e.g. the refinery's production capacity with its pre-conversion reformers, the production capacity following the installation of new HDS units, and the like) as either an increase in gasoline production volume or as an increase in the economic value of the fractions or products produced therefrom. The net capital cost for the steps of converting and replacing the at least one hydrocarbon reformer, in light of the increase in the net present value, is then greater than the net present value of the first production capacity minus a net capital cost for installing a new desulfurization system.
 In one embodiment, a continuous catalytic reformer replaces one or more of the old reformers in the refinery. The continuous catalytic reformer increases the percentage of high octane gasoline or blendable fraction of gasoline product produced per barrel in comparison to the percentage produced by the refinery using the original hydrocarbon reformer, if any, and a new desulphurization system. The fuel products produced by a refinery's desulfurization systems have a sulfur content which has been reduced to a level in compliance with a published government standard.
 In one embodiment, the increase in the percentage of high octane gasoline or fraction of gasoline per barrel of crude oil is greater than about 5% per barrel. Optionally, this increase in the percentage of high octane gasoline or fraction of gasoline produced per barrel can be between 5% to 27%.
 In one embodiment, at least two catalytic reformer systems are converted into desulfurization systems, and the net present value of the fuel produced by the refinery using the continuous catalytic reformer and the desulfurization systems, is greater than the net present value of the fuel produced by the refinery using at least two catalytic reformers prior to their conversion to desulfurization systems. In one embodiment, the desulfurization system reduces the sulfur content of the gasoline produced by the refinery to about 30 ppm or less. Optionally, the desulfurization system reduces the sulfur content of the diesel fuel produced by the refinery to about 15 ppm or less.
FIG. 1 depicts a simplified process flow diagram for the conversion of a BTX (benzene, toluene, xylene) reformer unit into a gasoline HDS unit.
FIG. 2 depicts a simplified process flow diagram for the conversion of a Motor reformer unit into a diesel fuel HDS unit.
FIG. 3 is a chart depicting projected net present value as a function of increasing diesel fuel and gasoline prices in cents per gallon.
 For the purposes of this disclosure, the terms “petroleum” and “crude oil” are interchangeably used to refer to the complex mixture of hydrocarbons obtained from oil fields. Depending on its geographic source, crude oil may contain about 35% to 65% (by volume) saturated hydrocarbons and olefins, 5% to 25% aromatics, 15% to 55% naphthenes, and about 0.2 to 3% sulfur (% wt).
 The terms “sulfur compounds” and “sulfur impurities” refer to the elemental sulfur and sulfur containing organic compounds found in crude oil and distillation fractions, which generally include, but are not limited to, hydrogen sulfide and sulfur containing hydrocarbons such as sulfides, thiols, and thiophenes.
 The terms “distillation” and “fractionation” are alternatively used to refer to the step in the refining process separating crude oil into various fractions. Each hydrocarbon fraction thus produced has its own characteristic boiling point range, and can be generally classified as low boiling point fractions, middle boiling point fractions, or high boiling point fractions. Low boiling point fractions generally include low boiling point hydrocarbons such as methane, ethane, and propane. Middle boiling point fractions generally include (or may be used to produce) naphthas, kerosene, benzene, toluene, naphthalene, and distillate fuels (i.e. diesel fuel, virgin fuel oil, and virgin heating oil). The high boiling point fractions include (or maybe used to produce) lubricating oils and tars, and the residue or bottom fraction contains tar and coke.
 The term “naphtha” is used to refer to the middle boiling range hydrocarbon fraction or fractions that are major components of gasoline, while the terms FCC naphtha and FCC gasoline refer to naphtha which has been produced by the process of fluid catalytic cracking.
 The terms “feed” and “feedstream” refer to the stream being sent to a reactor or treatment process while the terms “product stream” and “outlet stream” refer to the stream after it leaves the reactor. In a refinery, a hydrocarbon fraction may be referred to as both a feedstock and reaction product when between multiple steps in a treatment process. For example, the outlet stream of a diolefin treater may become the feed for blending gasoline.
 The term “reforming” refers to processes of thermally or catalytically converting hydrocarbon fractions into fractions having improved characteristics (e.g., a higher octane rating). Reforming encompasses processes that include cracking, polymerization, hydrotreating, dehydrogenation and isomerization reactions used to improve or convert hydrocarbon fractions into higher octane fractions. The term reformate refers to a hydrocarbon reaction product which has been subjected to one or more reforming steps.
 The term “cracking” refers to the process of breaking higher molecular weight (MW) hydrocarbon molecules into lower MW molecules, typically in the presence of a catalyst. In the refining process, cracking may be performed with reaction conditions including elevated temperatures, elevated pressure, the presence of a catalyst or a combination of these parameters.
 The term “desulfurization” refers to any of the several processes of chemically removing sulfur compounds from hydrocarbon fractions. Hydrodesulfurization is an example of these processes, as well as an example of catalytic hydrotreating. Hydrodesulfurization typically uses a catalyst and hydrogen gas to remove sulfur impurities by converting them to H2S and lower molecular weight compound. Hydrotreating processes may also be used to remove nitrogen containing compounds (e.g. amines, amides, etc.).
 The term “blending” as used herein, refers to the process of mixing different hydrocarbon fractions into a single product, typically the mixing of naphtha with reformates and raffinates or other hydrocarbon fractions to produce a gasoline product with a predetermined octane rating (an anti-knock rating system), vapor pressure, or other desirable characteristic. Blending may take place at the refinery, the gasoline retailer, or even at the gas pump at a gas station.
 The term “light straight run” (LSR) refers to the low to mid range boiling point hydrocarbon fractions taken directly from a side stream off the crude distillation column or downstream column if a series of distillation columns is used. An LSR also typically contains sulfur impurities.
 The term “net present value” (NPV) refers to the sum of the present values of future income (e.g. net income per year). To reduce a future income to a present value, a discount rate representing, inter alia, the lost opportunity costs of an investment, is applied to the future income according to the following formula:
NPV=I 0 +I n/(1+r)n
 Here, “I” represents the net yearly income obtained from the investment. “I” may also be negative, e.g. when capital expenditure, losses, or operating costs are greater than income for a given year. The “r” represents the discount rate, and the “n” represents the year of the income.
 In the present invention, an NPV was estimated for several refinery unit additions and modifications, with regard to meeting the upcoming LSG and USLD regulations, including the addition of new units, the conversion of existing units to new uses, and upgrading equipment where appropriate. The NPVs were calculated 15 year terms and including the net initial installation costs
 To effectively perform this analysis, several assumptions and exclusions were made where appropriate in order to ensure the consistency of the comparisons, and in order to estimate the costs and income associated with each unit addition or modification. These assumptions include, for the reform-to-HDS conversion examples, that a new continuous catalytic reformer is also installed to replace the converted reformers, with an estimated initial investment of approximately $160 million dollars based on past installation costs for comparable reformers. The assumptions also include that all existing equipment and infrastructure in the reformers being converted are in good and working condition. Also, the NPV estimates for converting the reformers were made based upon fiscal year 2001 budget prices, while future income estimates assumed a 20% tax rate as well as a 12% discount rate. It is understood that these and other assumptions/exclusions in the analysis may be made, as long as consistency is maintained between the comparisons, particularly the comparisons of the NPVs associated with converting existing units to the NPVs of installing new units.
 The main process operating conditions for both the LSG and ULSD HDS systems were also estimated in order to guide the selection of materials and equipment. The LSG HDS main process operating conditions, such as reactor temperatures, pressures, hydrogen purity, etc, were based on engineering studies performed for existing refineries, and on reviews of existing Mobil desulfurization technologies. Desulfurization technologies reviewed included
 Exxon Mobil SCANFining IFP and Prime G+. In short, the main process operating conditions were 350 psig (pounds per square inch gage) reactor pressure, 700° F. end-of-run temperature, 75% hydrogen purity, and 2.0 LHSV (liquid hourly space velocity).
 The ULSD HDS main process operating conditions were based primarily on studies published by the National Petroleum Conference (June 2000) and published data. In short, the main process operating conditions were 800 psig reactor pressure, 750° F. end of run temperature, 75% hydrogen purity, and 1.0 LHSV. For both the LSG and ULSD conversions, current equipment and piping materials in each case were evaluated for use with these process conditions.
 The capital cost estimates for converting the existing reformers in particular were derived from engineering studies which determined what equipment and material were required to convert two gasoline reformer units into desulfurization treatment units. The equipment and material included as part of the estimates for each example includes, for example, but without limitation, new or modified reactors, towers, vessels/drums, heaters, heat exchangers, pumps, and compressors. Flow diagrams for the converted, reformer-to-hydrodesulfurization units (HDS) are disclosed herein in FIG. 1 and FIG. 2.
 The conversion of existing reformer units into HDS units may be made with only minor modifications as described below. The HDS units generally follow typical HDS process flows. In these process flows, a feed is mixed with hydrogen gas (e.g., recycled hydrogen and make-up hydrogen.) The feed is then heated to reactor inlet conditions by heat exchange with the hot reactor outlet stream, by heat provided by a feed heater, or both. The feed then travels to one or more reactors, where desulfurization occurs. The now hot product stream (i.e. the outlet stream) is routed back to the heat exchanger in order to heat the reactor feed. An air cooler reduces the temperature of the reactor effluent, and the liquid and vapor phases are separated in a low temperature, high pressure separator. The liquid is sent to a stabilizer and the H2 vapor to an amine tower, which removes the H2S in the H2 vapor.
 The vapor can then be returned to the hydrogen recycle stream. Installation of a new amine scrubber is included in the following conversions, but it is recognized that some refineries may be able to use existing scrubber systems. Moreover, it is understood that alternative process flows may be used, in either the new HDS units or in the converted reformer HDS units, depending in part on the refineries existing equipment, available infrastructure or particular needs. It is also understood that the installation of metallurgical upgrades in the conversion of reformers to HDS units (hereinafter referred to as “a conversion”) can in some instances be delayed, thus reducing initial installation costs. However, increased monitoring would then be required to ensure that the higher sulfur content in the feed and increased process temperatures and pressures do not impair the function of the piping or equipment, or cause excessive corrosion thereof.
 Referring to FIG. 1, a BTX reformer with an approximate maximum capacity of 30,000 Bbl/day of light naphtha, dehexanizer bottoms and light unicrackate, was analyzed for conversion into an HDS unit running approximately 42,000 Bbl/day of FCC naphtha to produce LSG or high octane gasoline fractions (e.g. a high octane pool). This conversion requires several expansions and additions to the current unit, along with some material upgrades due to higher H2S concentrations in the reactor effluent stream. A summary of the costs for this conversion is listed in Table 1 below.
 The total estimated costs for this conversion is about $20 million. As seen in Table 1, the majority of the costs in Example 1 are due to new equipment and material upgrades. These material upgrades include relining reactors, material changes on the tube and shell side of heat exchangers, and piping changes to prevent or reduce corrosion.
FIG. 1 depicts a simplified process flow diagram depicting the placement of reactors, towers, heat exchangers, pumps and compressors in the converted unit, particularly those newly installed or requiring material upgrades for the conversion.
 Prior to conversion into a hydrotreating unit, the BTX reformer pretreated its liquid feedstock to remove sulfur. Before the liquid feedstock and hydrogen mix were fed into existing reactors they were treated at a hydro pretreater 30 for desulfurization. The pretreater reactor 30 was used to prevent reformer catalyst fouling by removing sulfur compounds. From the pretreater reactor 30, it was then fed to the BTX reformer reactors 34.
 The prior to conversion, the BTX reformer reactors 34 include four spherical reactors made of 1¼ Cr (chromium steel). The conversion entails, in this instance, lining the BTX reformer reactors with 321 stainless steel (ss). The piping between the reactors and hot feed/effluent exchangers are also lined or replaced with 321 ss.
 The BTX reformer's stabilizer tower 38 requires a partial retray for the design gas and liquid flows. The top 1′of the tower 38 preferably is lined with 316 ss, and contains a 347 ss distributor. The rest of tower 38 may be carbon steel (C Stl). The conversion in this instance also requires addition of a new amine scrubber tower 42., which also may be made of C Stl.
 The BTX unit includes a feed surge drum 46 and a stabilizer reflux drum 54. These drums will remain in the conversion. The conversion also includes modification of the effluent separator 50. Wash water may be needed to remove scale from the effluent coolers, thus a washer boot (not depicted) will also be installed in the conversion.
 The conversion also includes three new drums. Because of the addition of amine tower 42, the conversion includes a new carbon steel, vertical knockout drum (not depicted) with a 316 ss demister. The conversion also includes a new 230 ft2H2/Cl bed 58 to remove catalyst poisons from the make-up hydrogen. The third drum is degassing drum 62, which is fed by the amine scrubber tower 42 bottoms.
 The BTX reformer employs four heaters, but only two will be needed for the conversion to an HDS unit. Heaters 70, 74 supply a fired duty of 68 MMBtu/hr and 42 MMBtu/hr respectively. For the conversion, the first heater 70 is modified for use as a stabilizer reboiler heater, and its burners are modified for an absorbed duty of 28 MMBtu/hr.
 The second heater 74 may optionally be used as a reactor preheater, but a retube may be required due to the need for a material upgrade. In particular, higher sulfur content in the feed may require that the heater be retubed, and the piping between heater 74 and the reactors be replaced with 9 Cr.
 Heat Exchangers
 In the conversion, the four BTX cold feed/effluent exchangers 78 meet the duty requirements for the HDS unit, but the tubes should preferably be upgraded to 321 ss due to H2S in the effluent. Conversion of the hot feed/effluent exchangers 82 also preferably includes a tube replacement with 321 ss in the conversion, along with a shell upgrades to 321 ss. The duty requirements for the HDS units are not met with this set of the BTX units exchangers, so an additional exchanger (not depicted) is added in the conversion.
 Stabilizer feed/bottom exchanger 86 needs no modifications for the conversion, but may preferably be modified to provide additional surface area. An additional 10,000 ft2 can be added in parallel with exchanger 86 by adding three new exchangers (not depicted) made of C Stl. Two new exchangers will also be added to the stabilizer bottoms cooler 100 to provide an additional 9,600 Btu/hr. The stabilizer reboiler exchanger (not depicted) and the surface condenser (not depicted) are abandoned in the conversion, but the rest of the existing exchangers may remain in service without modification, including stabilizer condenser 94 and effluent condenser 98.
 The BTX unit's feed pump 104 has a rated capacity of 900 gpm (gallons per minute), but the additional volume provided to the HDS unit in the conversion will require 1260 gpm. A 500 HP motor is installed in the conversion of the reformer's feed pump to meet the HDS requirements. The BTX unit's stabilizer reflux pump 108 meets all the requirements for the HDS unit, and the surface condenser condensation pump (not depicted) is no longer needed. Two new pumps 112, 116 plus two backup pumps (not depicted) are also included in the conversion. A centrifugal pump 112 is added for the conversion of the stabilizer reboiler into a heater, and a piston pump 116 is added for the wash water.
 The conversion includes modifications to the recycle compressor 120. The reformer's recycle compressor 120 is a five stage centrifugal compressor, with a 10,000 HP motor. This compressor is larger than necessary for the HDS unit, and may be destaged to replace the steam driver with a 4000 HP motor.
 The BTX reformer of Example 1 was also analyzed for conversion into an HDS unit, processing an additional LSR feed. The feed in this example includes 42,000 Bbl/day of FCC naphtha and about 10,000 Bbl/day of LSR. However, a splitter unit (not depicted) is included in the conversion upstream from the HDS reactors, producing gasoline and fuel gas fractions.
 A 30/70 split was assumed for the feed, therefore, only 30,000 Bbl/day of FCC naphtha will be hydrotreated.
 A summary the costs for this conversion is contained in Table 2 below.
 The total cost estimate for Example 2 is $30 million dollars, a slightly higher figure than that for Example 1. As seen in Table 2, the additional feed requires the installation of a splitter and its ancillary equipment. The splitter costs amount to approximately 45% of the estimated cost, whereas the other 55% is due to modifications and some new equipment for the desulfurization portion of the unit. Here, the lower volume entering the HDS reactors reduced the number of modifications to be made in this portion of the unit.
 The reformer in this example was originally designed to process a maximum of 18,000 Bbl/day of heavy unicrackate and some light naphtha. The modified hydrotreater unit, will run 15,000 Bbl/day of VFO to be processed into ULSD. High reactor pressures and low space velocities will require expansions and additions to the current unit along with some material upgrades. The estimated cost for this conversion is $20 million dollars. As seen in Table 3 below, approximately half of the estimated costs for Example 3 stem from the installation of a new HDS reactor.
 The current reformer's reactors do not have the capacity for the HDS unit, and were not the proper material. Therefore, all of the original reactors were replaced with one new reactor meeting all the HDS unit criteria. Again, material upgrades are mainly due to the high pressure, high temperature hydrotreater loops and higher sulfur content in the feeds. The ULSD HDS temperature and pressure operating conditions are also higher than current conditions in the reformer; therefore, the conversion preferably includes the replacement of a number of metal components with higher grade metal.
FIG. 2 is a simplified process flow diagram depicting the placement of reactors, towers, heat exchangers, pumps and compressors, particularly those newly installed or requiring material upgrades.
 The Motor reformer unit has five small vertical reactors. In this instance, the reactors are too small to meet the needs for the HDS unit, and would all require a lining upgrade to 321 ss. Thus, in this example, it is more cost effective to replace all of the reformer's reactors with a new HDS reactor 124. The HDS reactor 124 has 2 beds, and a 347 ss distributor. Due to high sulfur content in the feed, the reactor will be made of 1¼ Cr and will be 321 ss lined. Preferably, 321 ss piping between the reactors and the feed/effluent exchangers should also be installed. Unlike Example 1, no SHU reactor is included for Example 3.
 The conversion includes no modifications to existing towers. However, an amine scrubber 128 is added as described above.
 All existing drums and vessels will be utilized for the conversion. No modifications are needed for the charge drum (not depicted), steam drum (not depicted), reactor product separator (not depicted), and the steam separator (not depicted). The conversion does include installing a water boot in the stabilizer overhead. Three new vessels and drums will be required as in Example 1, a compressor knockout drum (not depicted), and H2/Cl guard bed 136, and an amine degassing pot 140.
 The reformer has four heaters 146, but due to the exothermic nature of the HDS reaction, only one heater is used in the HDS system, providing a fired duty of 57,000 Btu/hr. The heater tubes may also be upgraded from 2¼ Cr, 1 Mo to 9 Cr.
 Heat Exchangers
 The conversion includes retrofitting the reformer's feed effluent exchangers 144, 148 with a new shell and tube made of 321 ss, since they are in the high pressure, high temperature hydrotreater loop. The rest of the reformer's heat exchangers may be used without modifications.
 The reformer's stabilizer pump (not depicted), overhead pump (not depicted) and the spare charge pump (not depicted) will need no modification for the conversion. However, the conversion does include replacing the motor of the charge pump 152 to increase the horsepower from 400 hp to 500 hp, due to the increased volume of the HDS feed relative to the reformer's feed.
 Examples 1 to 3 above assume that a modem reformer unit is built to replace the converted reformers. For example, a continuous catalytic reformer (CCR) with a capacity of about 45,000 Bbl/day, may be installed to replace both the Motor reformer and BTX reformer's discussed above. The CCR's annual operating cost is estimated to be about $10 million per year, and annual maintenance costs of about $3 million per year. The operating and maintenance costs of the CCR may be offset by a reduction in maintenance and operating costs of an existing H2 plant, which can be sold or taken off-line. Based on known CCR efficiencies, (capacity and yields), replacing the converted reformers with the CCR results in a $12 to $15 million/year increase in revenue. This increase in revenue stems from an estimated 4,500 Bbl/day increase in gasoline production, an increase in pool octane (the amount of high octane fraction available for blending into a gasoline fuel product) and, optionally, an increase in BTX (benzene, toluene, xylene) production. In essence, the modern reformer increases the potential production capacity of gasoline (per barrel) by more than about 3-5% per barrel. Depending in part on the efficiency and production capacity of the original informer this increase my be as much as between about 5% to 27% (per barrel). For the installation of a modern and highly efficient continuous catalytic reformer, the increase may be from 60-80% for the existing reformer to about 90-98% for the high efficiency catalytic cracking reformer. However, it is recognized that these incremental increases in efficiency may be based on the relative inefficiency of the reformer being converted and replaced.
 In the above examples, the summation of the savings, maintenance costs, operating costs, and improved gasoline production, results in an increase in yearly refinery income of approximately $12 million dollars. Taking taxes and depreciation into account, as well as the capital costs for converting the two reformers and installing the CCR, the NPV from Equation 1 is −88.3 at current gasoline and diesel fuel prices.
 The capital investment required to install and bring online new gasoline and diesel hydrodesulfurization systems is summarized briefly in Table 3 below.
 In essence, the initial capital cost for installing these two systems is approximately $135 million, expended over 4 years. In addition, starting from year one of operation, a new gasoline HDS unit costs roughly $8 million dollars per year to operate in order to maintain current production levels of gasoline (as low sulfur gasoline). A new diesel fuel HDS unit costs $7.5 million per year to operate. Once operational, an estimated $2 million would need to be spent on HDS unit maintenance (based on the historical maintenance costs of currently employed hydrotreating systems). These costs would be offset by a $5 million saving increase in revenue per year starting in 2006, as diesel fuel would no longer be treated by the FCC unit, thus creating downstream capacity and the potential for incremental crude throughput. Moreover, the HDS units do not produce an increase in gasoline production, pool octane, or BTX production. Taking taxes and depreciation into account as in Example 3, as well as the capital investment costs for installing new HDS units from Table 3, the installation results in a decrease in future income of about $12 million per year. The NPV from Equation 1 in this instance is −132.9, again, at current gasoline and diesel fuel prices.
 Thus, although the $135 million dollar initial capital investment for new HDS units does appear to compare favorably to the $200 million estimated for converting a BTX and Motor reformer into HDS systems and building a modem reformer unit, when the increased yearly income of the latter method is taken into account, the NPV of the latter method is greater than the NPV for the new HDS units.
FIG. 3 is a chart illustrating NPV as a function of an incremental change in the price of diesel fuel and gasoline. As seen in FIG. 3, a refinery implementing the inventive method would be able to have a zero net loss from implementing the new regulations and increasing the price of gasoline and diesel fuel by only about 1.0 to 1.9 cents per gallon. In contrast, installing HDS systems would require an increase of greater than 2 cents per gallon in order for the refinery to break even.
 The foregoing recitation of the invention is offered for the purposes of illustration only. It is recognized that the embodiments described herein may be modified or revised in various ways without departing from the spirit and scope of the invention. Instead, the scope of the invention is intended to be measured by the appended claims.