Environmental Protection Agency

Stancil & Co.____________________________________________________________________________________________________________________________________________

Peer Review of Refining Industry Cost Model

Report Date: September 27, 2013
Environmental Protection Agency

Peer Review of Refining Industry Cost Model

Table of Contents

Stancil & Co.____________________________________________________________________________________________________________________________________________

i

 Page

 Number

I. Introduction............................................................................................................................................................. I-1

II. Summary and Conclusions.................................................................................................................................... II-1

III. General Observations

Model Methodology ............................................................................................................................................ III-1

Model Structure and Capacity Assumptions........................................................................................................ III-1

FCC Naphtha Post-Treater Capacity Calculations .............................................................................................. III-2

Crude and Coker Naphtha Hydrotreater Capacity Calculations........................................................................... III-4

FCC Feed Sulfur Level Calculations ................................................................................................................... III-4

IV. EPA Inquiry A......................................................................................................................................................... IV-1

V. EPA Inquiry B......................................................................................................................................................... V-1

VI. EPA Inquiry C......................................................................................................................................................... VI-1

VII. EPA Inquiry D......................................................................................................................................................... VII-1

VIII. EPA Inquiry E......................................................................................................................................................... VIII-1

IX. EPA Inquiry F ......................................................................................................................................................... IX-1

Vendor 1............................................................................................................................................................. IX-1

Vendor 2............................................................................................................................................................. IX-3

Vendor 3............................................................................................................................................................. IX-3

X. EPA Inquiry G......................................................................................................................................................... X-1

XI. EPA Inquiry H......................................................................................................................................................... XI-1

XII. EPA Inquiry I .......................................................................................................................................................... XII-1

XIII. EPA Inquiry J.......................................................................................................................................................... XIII-1
Environmental Protection Agency

Peer Review of Refining Industry Cost Model

I. Introduction

Stancil & Co.____________________________________________________________________________________________________________________________________________

I-1

The Environmental Protection Agency (EPA) Office of Transportation and Air Quality (OTAQ) contracted with ICF International 

to provide three independent peer reviews to separately assess a Refining Industry Cost Model (EPA Model) that was created 

to assist in finalizing Tier 3 regulations that affect gasoline sulfur standards. Stancil & Co. (Stancil) was one of the three 

subcontractors chosen to perform one of these peer reviews. 

The EPA model was originally developed by APT (Mathpro) for the EPA to estimate the cost of benzene control under Mobile 

Source Air Toxics 2 (MSAT2). The model was later revised by the EPA OTAQ to include representations for gasoline 

desulfurization along with the associated capital and operating costs. The EPA model is a refinery-by-refinery representation 

of all United States (U.S.) refineries that calculates the cost for each refinery to reduce gasoline sulfur levels to the standards 

proposed by the Tier 3 regulations.

Stancil’s review focused on model methodology, the various assumptions and data used, how data was interpreted and 

incorporated into the model, and the logic equations used throughout the model. Our goal was to assess the EPA model’s 

ability to determine refining costs for reducing gasoline sulfur levels from current levels to proposed Tier 3 levels and to 

recommend changes to improve model accuracy.

In addition to a general review of the model, the EPA requested specific review of 10 different areas of the model. Each of 

these areas will be addressed in separate sections of this report.
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Peer Review of Refining Industry Cost Model

II. Summary and Conclusions

Stancil & Co.____________________________________________________________________________________________________________________________________________

II-1

Stancil spent seven days with access to the Environmental Protection Agency (EPA) model to review the assumptions and 

logic that were used to calculate the refinery-by-refinery estimated costs for implementation of the proposed Tier 3 

regulations. We found that in order to understand the cost estimations for expansion of gasoline treatment capacity, a 

detailed understanding of the assumptions used for unit capacities and calculation of gasoline production volumes was 

necessary to understand the logic behind the technology choices made. The flow of discussion in this report will touch upon 

some issues with model structure, unit capacity assumptions, and fluid catalytic cracker (FCC) feed sulfur calculations in 

Section III General Observations and then proceed to discussions on the ten areas of interest identified by the EPA in 

Sections IV through XIII. 

By its nature, the EPA model is a very large and complicated model that attempts to model all individual United States (U.S.) 

refineries within a single Excel workbook. The number of logical statements that are used to handle all of the various refinery 

situations is also very cumbersome. The addition of actual process unit throughputs and actual finished gasoline volumes to 

the model in its latest incarnation added an additional level of complexity to the model. We agree there is benefit to validating 

the model’s accuracy by comparing the model theoretical yield predictions to actual volumes produced at the refineries; 

however, trying to do both tasks within a single model can be very confusing. We would recommend that the EPA consider 

doing the actual capacity and production balances in a separate workbook from the theoretical calculations and then compare 

the results of each model side-by-side. In our linear programming (LP) validation work, we find this side-by-side comparative 

analysis very helpful in improving model accuracy.

The addition of the actual unit throughput information into the model appears to have clouded the distinctions between the 

design capacities of the refinery process units and the operable capacities of the process units in the model. We are of the

opinion that design capacities reported by the Energy Information Agency (EIA) in terms of barrels per stream day (B/SD) are 

the appropriate benchmark capacity from which to size additional gasoline treatment equipment needed for Tier 3 regulations. 

The EPA model uses the lower annual averaged capacity reported by the EIA in terms of barrels per calendar day (B/CD) in 

its cost calculations. We believe that the “maximum FCC gasoline” volume should be calculated using EIA B/SD FCC 

capacities for sizing the new post-treaters rather than the current method of using EIA B/CD capacities. The same B/SD 

methodology should also be used in determining the size of additional naphtha hydrotreating requirements.

We believe that the EPA model understates the amount of sulfur in FCC feedstock for some refineries. The level of sulfur in 

FCC feed is used to estimate FCC naphtha sulfur concentrations that are then later used to determine some of the technology 

choices in the EPA model. This adjustment may affect only a few refineries, but we believe it is one of the assumptions that

are important to recognize. Part of the downside to using a refinery-by-refinery model is recognizing all of the smaller effects 

that can add to refiners’ costs.
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Peer Review of Refining Industry Cost Model

II. Summary and Conclusions

Stancil & Co.____________________________________________________________________________________________________________________________________________

II-2

In our opinion, light straight run (LSR) gasoline volumes appear to be underestimated in the EPA model. This conclusion 

stems from two areas: the first being the correlation equation used to calculate the LSR yield from crude oil appears to 

underestimate LSR yield based on some of the test cases we assessed. The second area is the cut point between LSR and 

heavy straight run (HSR) naphtha used in the EPA model is lower than what we routinely see at the refineries we work with. 

Directionally, using a higher LSR cut point will help with the model’s overproduction of reformer feed. Additional LSR volume 

may lead to some additional naphtha treating requirements at some refineries.

In general, we believe that assuming the reformate yields at 90 research octane number (RON) currently used in the EPA 

model may be overestimating the volume of reformate production. While we have no data to show what the average might be 

for the entire U.S., the refineries we work with on a routine basis typically operate in a range of 94 RON to 98 RON for semiregenerative reformers and up to 100 RON for CCR units depending on their specific octane situation.

A lot of effort is spent in the EPA model to balance theoretical yields of gasoline against actual production volumes. A number 

of corrections are used to shift gasoline into distillates. Some of these are undercutting reformer feed into jet fuel, 

undercutting FCC naphtha into light cycle oil (LCO), and shifting hydrocracker (HDC) operations from naphtha mode into 

diesel mode. All of these are legitimate means that refiners use to shift gasoline production to jet fuel or diesel production. In 

the model, these shifts are done as needed to force the gasoline and reformer feed to balance with actual throughput and 

production values for each refinery though there is no data available to verify that a particular refiner is actually operating the 

way that the model assumes. One way the EPA could lend credence to the methodology would be to perform a distillate 

balance to see if the theoretical yields match actual data. 

Based on the comments of previous peer reviews and our review of the model, it appeared to us that most of the vendor 

technical information on FCC naphtha post-treater capital and operating costs were unchanged since the last peer review. 

Our opinion of the minimum investment cases provided by Vendor 1 for evaluating FCC naphtha post-treater costs for

reducing FCC naphtha sulfur concentrations from 75 ppm to 25 ppm is they do not seem to be practical from an operations 

standpoint of avoiding periodic FCC shutdowns or rate reductions due to shorter post-treater catalyst run cycles. While we 

understand that there may be some potentially new developments in technology and catalyst design that may improve the 

performance of existing FCC naphtha hydrotreaters, there was no information available in Vendor 1’s information package to 

evaluate the basis of these claims. 

We noted that the EPA incorporated some of this data in the refining cost model, but we were unable to understand the EPA’s 

logic in how investment costs were determined for the 200 ppm and 800 ppm FCC naphtha sulfur cases in the model. We 

noted that the previous peer reviewers expressed similar concerns with the same set of data in their 2011 report. If the EPA

plans to use this minimum investment data, we would recommend that the additional costs associated with FCC throughput 

reductions or shutdowns that were recommended by the vendor be included in the EPA model analysis.
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Peer Review of Refining Industry Cost Model

II. Summary and Conclusions

Stancil & Co.____________________________________________________________________________________________________________________________________________

II-3

We were able to validate Vendor 3’s cost estimate for building a two-stage, post-treating unit based on the construction costs 

we have seen working with various clients and construction companies. Including the 30% contingency factor recommended 

by the vendor, Vendor 3’s cost estimates for building a two-stage FCC naphtha hydrotreater match fairly well with what our 

own cost curves would predict. While we do not have any cost curves for adding a second-stage reactor alone, the fact that 

we can match the vendor’s capital cost for a full two-stage unit allows us to have confidence in the method used by the EPA 

for estimating Vendor 3’s second–stage only costs. 

We would agree that extractive treating of butane is widely practiced today. This does not, however, necessarily assure that

an ultra low sulfur butane product is being uniformly produced. We have noted sulfur amounts in blending butane that range 

anywhere from 0 ppm to 30 ppm at the refineries we work with. Many refiners store excess butane production off-site during 

the summer months when gasoline Reid vapor pressure (RVP) specifications are low and bring them back to the refinery for 

gasoline blending in the winter months when RVP specifications are higher. Specifications for refinery-grade butane at some 

off-site storage facilities allow up to 140 ppm total sulfur. Even if refiners are able to produce ultra low sulfur butane for 

gasoline blending, they may need to incur extra expense to set up dedicated butane storage facilities or potentially add more

caustic treating capacity to treat any butane being brought back from storage to suitable sulfur levels for Tier 3 gasoline 

blending. 

Based on assay data that shows sulfur content in LSR gasoline is not correlated with the amount of sulfur in crude oil, we 

would have to conclude that there is not particular crude sulfur percentage cut-off point where extractive caustic treatment of 

LSR would stop and hydrotreating would begin. The cut-off point in the EPA model is currently set at a crude sulfur content of 

1.0 Wt.%. One possible issue with extractive caustic treatment of gasoline boiling range material, such as LSR, is that sulfur 

removal efficiency is lower than for lighter LPG feedstock. We have only anecdotal evidence from a competing vendor that 

this efficiency may be as low as 90%. At this efficiency level, LSR feed with sulfur concentrations over 10 ppm could become 

problematic if a refinery were depending on a treated LSR sulfur content of 1 ppm for gasoline blending as in the EPA model. 

We would recommend that the EPA try to verify from the vendors, what is the sulfur removal efficiency for extractive caustic 

treatment of LSR materials.
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Peer Review of Refining Industry Cost Model

III. General Observations

Stancil & Co.____________________________________________________________________________________________________________________________________________

 III-1

Model Methodology

In reviewing the EPA model, we noted that while the model’s primary function is to estimate the refinery-by-refinery costs of 

proposed lower Tier 3 gasoline sulfur standards, a considerable amount of the model is devoted to estimating the refinery-byrefinery volume and sulfur level of gasoline blendstocks produced. The volume of blendstocks produced at each refinery was 

initially estimated based on process unit capacity information and unit yield assumptions obtained from various sources. The 

total volume of the estimated blendstocks was then compared to the total volume of the actual gasoline production from each 

refinery. For refineries where the estimated blendstock volumes and/or total gasoline volume did not match actual values, a 

number of options are built in to adjust the model’s yield assumptions to bring the estimated and actual volumes into balance. 

Several of these adjustment options are discussed in more detail later in this report.

Sulfur concentrations for each of the gasoline blendstocks were estimated from various assumptions provided by literature 

sources, refinery consultants, and technology providers. FCC naphtha is assumed to be the primary contributor to the amount 

of sulfur in the gasoline pool, but the concentrations vary widely from refinery-to-refinery. To account for this variability, the 

EPA model calculates what the FCC naphtha sulfur concentration would need to be given the volume of other gasoline 

blendstocks and their sulfur concentration to produce the sulfur concentration in the total volume of finished gasoline. The 

model uses these baseline volume and sulfur assumptions to determine the amount of sulfur removal needed to meet the 

proposed Tier 3 standards. 

The model has incorporated various equipment configuration options based on data provided by literature sources, refinery 

consultants, and technology providers to remove the amount of sulfur needed. This data includes estimates for capital and 

operating costs needed to assess the economic impact of the Tier 3 regulation. 

Model Structure and Capacity Assumptions

Review of the EPA model found that process unit capacity data used in various formulas throughout the model referenced at 

least four different worksheets. Capacity data from the Energy Information Administration (EIA), the Oil & Gas Journal 

(O&GJ), and the Office of Air Quality Planning and Standards (OAQPS) were used. 

Initially, many of the capacities appeared to be mislabeled as being B/SD capacities when they were actually B/CD capacities. 

Actual throughput volumes from the OAQPS, which are essentially B/CD capacities for the year they occurred, were not 

designated as being actual rates and were often mislabeled as being B/SD capacities as well. After looking at how the 

capacities data was being used, we came to the conclusion that the EPA was using the actual rates from OAQPS to be B/CD 

capacities and the B/CD rates reported in the annual EIA capacity report were being used as B/SD capacities.
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Peer Review of Refining Industry Cost Model

III. General Observations

Stancil & Co.____________________________________________________________________________________________________________________________________________

 III-2

EIA reports all unit capacities as B/SD in their annual refining capacity report. Some of the primary operating units 

(atmospheric crude distillation, coking, reforming, hydrocracker (HDC), and fluid catalytic cracking (FCC)) are also reported in 

B/CD. Refiners are required to report these capacities to the EIA annually on form EIA-820. The B/SD capacity of a unit 

represents the engineering design capacity a unit is capable of when operating at over a 24-hour period. Usually, these are 

capacities that have been demonstrated during past operating periods and define the benchmark flow rates for new 

equipment design considerations. Downstream conversion units such as the coker, HDC, and FCC, will routinely operate at 

their B/SD capacities for periods of time until some operating or maintenance problem forces the unit to slow down. Since the 

amount of crude a refinery can process is typically limited by the capacity of these downstream conversion units, there is 

usually an economic incentive to maximize their throughput. 

The B/CD capacity of a unit represents the total rated capacity and is the amount of input that a unit can process under usual 

operating conditions during a year. The total amount of input reflects throughput reductions resulting from various limitations 

that can be expected to occur throughout the year. The average amount of this annual throughput is expressed in terms of 

capacity during a 24-hour period. The B/CD capacities refiners report will usually take into account anticipated limitations 

arising from turnarounds, equipment inspections, routine maintenance and repairs, environmental constraints, as well as 

types and grades of feedstock inputs and product produced. Barrel per calendar day capacities for the atmospheric crude unit

may also include the limitations of these downstream units.

Actual annual average operating capacities, like those reported by the OAQPS, will reflect all of the above limitations in 

addition to unexpected unit downtimes or rate curtailments that result from weather events, accidents, fires, inventory 

containment, or economic conditions. The capacity utilization percentages reported by the EIA represent these actual 

average capacities divided by the B/CD rates. It is not unusual to see capacity utilization percentages of over 100% reported 

if a refinery has experienced a period of good operations.

FCC Naphtha Post-Treater Capacity Calculations

One area we have an issue with regarding the use of B/CD and/or B/SD unit rates in the EPA model is with the calculations to 

determine capacity requirements for FCC naphtha post-treating units that are used to determine capital costs. For the 

calculation of maximum FCC naphtha volume, the EPA model uses the EIA B/CD capacity for the FCC unit multiplied by the 

PADD-by-PADD naphtha yields in the “process inputs” sheet. This B/CD volume of FCC naphtha, denoted as “max FCC 

gasoline” in the EPA model later becomes the design basis for determining the FCC naphtha treating costs associated with 

going to a 10 ppm or 5 parts per million (ppm) gasoline sulfur standard. In our opinion, the “max FCC gasoline” volume 

should be calculated based on the EIA B/SD capacities for the FCC rates.
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Peer Review of Refining Industry Cost Model

III. General Observations

Stancil & Co.____________________________________________________________________________________________________________________________________________

 III-3

As mentioned previously, the B/SD capacity of a unit represents the engineering design rate a unit is capable of when 

operating, and downstream processing units like the FCC routinely operate at their B/SD capacity over various periods of 

time. One of the principals in designing refinery equipment is to ensure that the new equipment is sized such that it will not 

become a limiting factor in any foreseeable mode of operation. In terms of the “max FCC gasoline” volume calculation, this 

would include not only the gasoline volume associated with the FCC B/SD capacity, but also with the unit operating in a 

maximum gasoline mode. To do otherwise, the refiner would have to make an economic decision to purposefully derate the 

unit’s capacity. If the FCC naphtha post-treater size were designed on the B/CD capacity basis, it would not have the spare 

make-up capacity to get back to average in the event the unit had a planned or unplanned downtime. 

We noted in the EPA model that an average overdesign factor of 7.5% was applied to the “maximum FCC gasoline” volume 

calculated from the B/CD FCC capacities based on a 5% to 10% over-design range used by one of the vendors in their FCC 

naphtha post-treater cost calculations. Our understanding of using an over-design factor in equipment design is to allow for 

fluctuations in service that would be expected or anticipated with respect to normal operation. For instance, if an FCC were

operating at its design capacity, naphtha produced while the unit was in a normal steady-state operation would be at some 

constant design flow rate. Occasionally, equipment malfunctions or perhaps off-specification feedstock can cause what is 

called a unit “upset” making the production yields change rapidly and cause flow rates to vacillate up and down for a period of 

time until the unit controls can return the unit to steady-state operation. The over-design factor allows for enough extra 

equipment size to handle the flow rate during the “up” portion of the vacillation to prevent overloading.

Other purposes for using an over-design factor are to allow for decreases in equipment and catalyst efficiency over a period of 

time between turnarounds and to provide allowances for future developments and expansion. We do not believe the overdesign factor is intended to bridge the gap between B/SD and B/CD capacities. 

In summary, it is our opinion that maximum FCC naphtha rates used to calculate capital costs for installing FCC naphtha posttreaters in the EPA model should be based on the B/SD FCC capacities shown in the EIA annual capacity report. We also 

believe that the maximum FCC naphtha rates should be for a full boiling range product that would include the 350°F to 400°F 

fraction as seasonal demands for gasoline can often drive refineries into maximum gasoline production mode. Calculation of 

FCC naphtha post-treater operating cost for utilities, hydrogen consumption, and octane loss on a B/CD basis, as they 

currently are in the EPA model, are appropriate for determining ongoing annual costs. Our primary concern is to be sure all of 

the one-time capital outlays refiners will face are captured. 
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Peer Review of Refining Industry Cost Model

III. General Observations

Stancil & Co.____________________________________________________________________________________________________________________________________________

 III-4

Crude and Coker Naphtha Hydrotreater Capacity Calculations

We have a similar issue with the use of B/CD and/or B/SD unit rates in the EPA model for the calculations to determine 

capacity requirements for crude and coker naphtha hydrotreating units that are used to determine capital costs. As with the 

“maximum FCC naphtha” volume calculation, maximum crude and coker naphtha volume calculations also use B/CD 

capacities to determine the amount of naphtha hydrotreater expansions needed for Tier 3 regulations. Other issues related to 

the volume of naphtha produced from crude are discussed in more detail in Section IV, EPA Inquiry A and Section VI, EPA 

Inquiry C. We believe the methodology for determining naphtha hydrotreater capacity requirements to meet Tier 3 standards 

should be done in a similar manner as that for FCC naphtha post-treater calculations. 

FCC Feed Sulfur Level Calculations

We are in agreement with the way the EPA model calculates FCC feed sulfur concentrations for virgin vacuum gas oil (VGO) 

produced from crude oil. The graph below shows the relationship of crude VGO sulfur to crude sulfur used in the EPA model.

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Virgin FCC Feed Sulfur, Wt.%

Crude Sulfur, Wt.%

EPA Model - Crude VGO Sulfur Versus Crude 

Sulfur
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III. General Observations

Stancil & Co.____________________________________________________________________________________________________________________________________________

 III-5

Data from Stancil’s crude assay database is shown on the graph below. While this graph represents more crude types with a 

wider range of sulfur concentrations, the general relationship yields crude VGO sulfur concentrations similar to the range of

crude sulfur samples used in the EPA model.

There are, however, other gas oil streams produced in some refineries that are processed in the FCC unit that can add 

significantly to the sulfur load. These would include heavy gas oils produced from coker units, deasphalted oils produced from 

solvent deasphalting units, atmospheric residuum, and vacuum residuum.

The most common stream is heavy gas oil produced from coker units. Coker heavy gas oil (HGO) yield is typically around 

25 volume percent (Vol.%) of the vacuum residuum feed to a coker unit. The sulfur concentration of the coker HGO can 

roughly be estimated as about 1.1 times the sulfur concentration of the vacuum resid. The graph shown below represents the 

relationship of vacuum residuum sulfur to crude sulfur based on a vacuum residuum cut point of 1,000°F. 

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Virgin FCC Feed Sulfur, Wt.%

Crude Sulfur, Wt.%

Stancil - Crude VGO Sulfur Versus Crude Sulfur
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III. General Observations

Stancil & Co.____________________________________________________________________________________________________________________________________________

 III-6

Assuming a crude sulfur of 2.0 weight percent (Wt.%), vacuum residuum would contain about 4.0 Wt.% or 40,000 ppm sulfur. 

Coker HGO produced from this material would contain approximately 44,000 ppm sulfur. By comparison, virgin VGO 

produced from this same crude would have about 20,000 ppm sulfur. If the coker HGO comprised 10 Vol.% of FCC feed, the 

sulfur concentration would increase to 22,400 ppm, about 12% higher than a 100% virgin VGO feed. 

Since FCC feed sulfur content is used to estimate FCC naphtha sulfur calculations in the EPA model that affect some of the 

logic used to determine post-treating technology choices, we think at a minimum the coker HGO should be included in the 

FCC feed sulfur calculations for those refineries with coker units. 

y = -0.2093x2 + 2.3238x + 0.1941

R2 = 0.9316

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Vacuum Residuum Sulfur, Wt.%

Crude Sulfur, Wt.%

Stancil - Vacuum Residuum Sulfur Versus Crude 

Sulfur
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Peer Review of Refining Industry Cost Model

IV. EPA Inquiry A

Stancil & Co.____________________________________________________________________________________________________________________________________________

 IV-1

Inquiry A – Review the methodology for estimating the volume of light and heavy straight run naphtha which is based on a 

regression analysis of the API gravity and light straight run fraction from the assays of [13] crude oils. (This replaced the 

previous method of relying on similar correlation for the average quality of crude oil refined in each PADD).

Initially, one concern we had for the methodology being used in the EPA model to determine crude naphtha yields is that the 

13 crude oil assays used do not provide a large enough sample size to adequately develop generalized yield correlations with 

crude density in degrees American Petroleum Institute (°API) gravity for the various crude cut points in question. However, as 

Stancil will demonstrate with some examples, even with a larger sample size we are not convinced that the use of generalized 

yield equations will give accurate results for mixtures of various crudes types. 

While the 13 crude assays used in the EPA model provided some coverage of the crude types processed throughout the U.S. 

Stancil felt that a number of important crude types were left out of the mix. We did not see assays for shale oil (such as 

Bakken, Eagle Ford, etc.), heavy Canadian diluted bitumen crudes (such as Western Canadian Select (WCS), Cold Lake, 

etc.), or U.S. Gulf Coast (USGC) offshore crudes (such as Light Louisiana Sweet (LLS), Mars, Thunderhorse, etc.) 

represented. These crudes have been and will continue to comprise an increasing amount of the U.S. crude supply in the 

future. Based on our conversation during the kick-off meeting, it was our understanding that the number of crude assays 

available to the EPA for this study was limited, so we took the liberty of utilizing our crude assay database to develop some 

graphics to illustrate some of our concerns.

Using the yield data from the 13 crude oils in the EPA model, we graphically recreated the yield representations for the crude 

light straight run (LSR) gasoline and heavy straight run (HSR) naphtha to visually illustrate the distribution of data. Each 

graph shows the linear equation that was developed from the data to estimate LSR and HSR yields as a function of °API

gravity. We also developed similar graphs for LSR and HSR from a database of crude assays we maintain at Stancil. This 

data set contained approximately 175 different crudes that were recently updated. We recut the assays to the same 

distillation cut point temperatures as those in the EPA refining model. 

We note from the graph on the following page that the linear LSR yield correlation developed from the 13 crudes in the EPA 

model has a low coefficient of determination (denoted by R2) of 0.38 that is an indication the equation may have limited utility 

for producing precise results. Visually examining the data, one can conclude that the equation may produce reasonable yield 

results for crudes around 30°API (the median °API and LSR yield for all 13 crudes is 29.4 and 0.0506, respectively), but it is 

less clear how well it would predict LSR yields for crudes at 20°API or 40°API. 
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IV. EPA Inquiry A

Stancil & Co.____________________________________________________________________________________________________________________________________________

 IV-2

The following graph shows the LSR yield versus °API from the Stancil crude assay database mentioned previously. The best 

correlation for the data was a polynomial regression (R2= 0.82) that defines an upwardly sloping curve, which becomes more 

pronounced for crudes with gravities higher than about 35°API. We think it is important to consider the LSR yield relationships 

above 40°API as most shale oil currently being produced has gravities in the 40°API to 60°API range. Condensates and 

natural gasoline resulting from natural gas production often have gravities in the 60°API to 70°API range.

LSR Yld = 0.0021*API - 0.0125

R2 = 0.3760

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LSR Yield, Volume Fraction

 API Gravity

EPA Model - LSR Yields (C5-160 F)
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IV. EPA Inquiry A

Stancil & Co.____________________________________________________________________________________________________________________________________________

 IV-3

Another part of this graph we would like to point out is in the gravity region around 20°API by noting there are several outlying 

data points with high LSR contents of 6 to 11 volume percent (Vol.%). These outliers are the heavy Canadian bitumen crudes 

that are diluted with natural gasoline, condensates, and light naphtha for transportation and handling requirements. 

The graph below is the linear HSR yield correlation developed from the 13 crudes in the EPA refining model. We note that the 

correlation has a reasonably good R2 value and visual examination indicates the equation has a reasonably good fit for the 

data provided.

y = 2E-06x3- 8E-05x2 + 0.0022x + 0.0001

R2 = 0.8197

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LSR Yield, Volume Fraction

 API Gravity

Stancil - LSR Yield (C5-160 F)
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IV. EPA Inquiry A

Stancil & Co.____________________________________________________________________________________________________________________________________________

 IV-4

The following graph shows the HSR yield versus °API from the Stancil crude assay database mentioned previously. The best 

correlation for the HSR data was a polynomial regression with an R2 = 0.89 that defines an upwardly sloping curve similar to 

LSR, though less pronounced. As with LSR, we would reiterate the importance of considering the yield relationships above 

40°API to account for shale oil, natural gasoline, and condensate production.

HSR Yld = 0.0090*API - 0.1063

R2 = 0.8827

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HSR Yield, Volume Fraction

 API Gravity

EPA Model - HSR Yields (160 F-350 F)
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IV. EPA Inquiry A

Stancil & Co.____________________________________________________________________________________________________________________________________________

 IV-5

All of the preceding graphs have shown that there can be a considerable difference in the yields of LSR and HSR product 

from crudes of similar °API gravity. If one assumed that all of the crude oils a refinery processed were of the same °API 

gravity, using these generalized correlations can probably be justified as they provide an average yield for that particular set of 

crude oils. In practice, however, most refineries process a number of different crude oils with a wide range of °API gravities. 

A large part of the crude oil optimization process is to tailor a crude oil slate, so that the resulting yields will provide the best 

balance of feedstocks to optimally fill all of the downstream processing units. For example, a refinery with a coker unit will 

input a certain amount of low °API gravity crude to load that unit, another amount of medium °API gravity crude to load the 

FCC, HDC, and diesel hydrotreating units, and a balance of high gravity °API gravity crude to load the isomerization and 

reforming units. Ideally, the resulting crude mix will load the process units without producing excess feedstocks that might 

have to be sold. 

To test the various crude yield correlations for accuracy in predicting LSR and HSR yields with mixtures of crude oils, we put 

together some theoretical cases with typical crude types usually processed in USGC refineries. One set of crude slates 

resulted in weighted average crude °API gravities of 28.8°API, close to what Texas Gulf Coast refineries averaged in 2011. A 

second set of cases reflect what a crude slate might look like in 2017 as more foreign crude oils get replaced with domestic y = 9E-05x2 + 0.0021x - 0.0021

R2 = 0.887

0.00

0.10

0.20

0.30

0.40

0.50

0.60

0.70

0.0 10.0 20.0 30.0 40.0 50.0 60.0 70.0 80.0

HSR Yield, Volume Fraction

 API Gravity

Stancil - HSR Yield (160 F-350 F)
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Stancil & Co.____________________________________________________________________________________________________________________________________________

 IV-6

and Canadian grades. We should emphasize that these are not actual crude slates and are being used only for the purpose 

of assessing yield correlation accuracy.

The first case, summarized below, uses crude data from the EPA refining model to calculate weighted averages for °API, LSR 

yield, and HSR yield based on the indicated percentages of Maya, Saudi Medium, Bonny Light, and West Texas Intermediate 

(WTI) crudes. We then used the EPA model equations to calculate LSR and HSR yields based on the weighted average 

crude °API. For this case, the EPA model equations under predicted both LSR and HSR by about 1.0% and 0.8%, 

respectively. 

EPA Data - Conventional Crude 2011

% of Crude Vol. Fraction

Crude Slate °API LSR HSR

Maya 45 22.3 0.0514 0.1252

Saudi Medium 25 30.4 0.0527 0.1504

Bonny Light 15 34.6 0.0389 0.1880

WTI 15 39.7 0.1028 0.2592

Weighted Average 100 28.8 0.0576 0.1610

EPA Equation Results 28.8 0.0479 0.1527

The second case with crude data from the EPA refining model is summarized below. For this case, we assumed Saudi 

Medium and Bonny Light are replaced with a mixture of Bow River and additional WTI. We allow a higher weighted average 

°API to reflect the abundance of light crude production expected in the future. For this case, the EPA model equations under 

predicted both LSR and HSR by about 2.7% and 1.4%, respectively. 
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 IV-7

EPA Data - Conventional Crude 2017

% of Crude Vol. Fraction

Crude Slate °API LSR HSR

Maya 35 22.3 0.0514 0.1252

Bow River 20 23.4 0.0719 0.1029

Bonny Light 0 34.6 0.0389 0.1880

WTI 45 39.7 0.1028 0.2592

Weighted Average 100 30.4 0.0786 0.1810

EPA Equation Results 30.4 0.0512 0.1669

The next set of cases used a number of different crude oils in the Stancil crude assay database to simulate the 2011 case and 

the Stancil LSR and HSR correlation equations discussed previously to estimate the LSR and HSR yields. For this case, a 

number of other crudes typically processed on the USGC (Bachaquero 17 (BCF-17), Mars, LLS, and Saharan Blend) were 

included. For this case, the Stancil equations prediction for LSR was very close, but the HSR prediction was low by about 

2.9%.

Stancil Data - Conventional Crude 2011

% of Crude Vol. Fraction

Crude Slate °API LSR HSR

BCF-17 15 16.9 0.0136 0.0456

Maya 20 21.6 0.0242 0.1292

Mars 10 29.0 0.0593 0.1521

Saudi Medium 22 30.8 0.0464 0.1833

Bonny Light 18 35.3 0.0543 0.2332

LLS 10 36.5 0.0550 0.1854

Saharan Blend 5 45.5 0.1309 0.2691

Weighted Average 100 28.8 0.0448 0.1622

Stancil Equation Results 28.8 0.0449 0.1331
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Stancil & Co.____________________________________________________________________________________________________________________________________________

 IV-8

The final case with crude data from the Stancil database is summarized below. For this case, we assumed Saudi Medium, 

Saharan Blend and some BCF-17 are replaced with a mixture of WCS (diluted bitumen), WTI, and Eagle Ford Light (shale 

oil). As in the other future case, we allow a higher weighted average °API to reflect the abundance of light crude production 

expected in the future. For this case, the Stancil equations under predicted both LSR and HSR by about 0.6% and 2.9%, 

respectively. 

Stancil Data - With Unconventional Crude 

2017

% of Crude Vol. Fraction

Crude Slate °API LSR HSR

BCF-17 10 16.9 0.0136 0.0456

Western Canadian Select 20 20.3 0.0624 0.0816

Maya 20 21.6 0.0242 0.1292

Mars 10 29.0 0.0593 0.1521

LLS 10 36.5 0.0550 0.1854

WTI 15 39.9 0.0718 0.2691

Eagle Ford Lt Shale 15 51.6 0.0917 0.3533

Weighted Average 100 30.3 0.0546 0.1738

Stancil Equation Results 30.3 0.0491 0.1445

Based on the information just presented, it is difficult to draw definitive conclusions for using the regression correlation method 

to estimate crude yields. For the case studies, all of the correlation equations tended to under predict LSR and HSR yields. 

Given the poor to mediocre R

2

values for the correlation equations, we would expect there would be some fairly large margins 

of error in the predictions. The other trend of note from the case studies is the LSR and HSR prediction error tended to be 

larger for the future cases, which contained higher percentages of light crude and unconventional crudes. For the purposes of 

determining Tier 3 costs, under prediction of LSR and HSR could have the effect of underestimating the costs required for 

naphtha hydrotreating projects if more LSR needs to be hydrotreated. We would recommend that the EPA consider 

performing a sensitivity case assuming a 1% higher LSR yield to see what the impact would be on Tier 3 costs. 
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V. EPA Inquiry B

Stancil & Co.____________________________________________________________________________________________________________________________________________

V-1

Inquiry B – Review the methodology of basing the refinery blendstock volumes for the reformer, alkylation unit, isomerization 

unit, aromatics unit and naphtha hydrotreater on actual throughput volume data from the Office of Air Quality Planning and 

Standards (OAQPS).

Alkylation unit throughput is reported as actual alkylate yield; therefore, using the actual alkylation throughput volumes from 

OAQPS will accurately reflect the blendstock volume in the gasoline pool. In the isomerization process, some hydrocracking 

reactions occur that result in some of the gasoline feedstock being converted to light gases. In our modeling work, Stancil

generally assumes the isomerate blendstock yield is about 98.5 Vol.% of the unit throughput. 

The volume of reformate blendstock produced at the reformer will vary depending on feedstock quality, operating pressure, 

and the octane severity the unit is operating. The data shown in the “process inputs” sheet of the EPA model appears to 

indicate an operating severity of 90 research octane number (RON) across every PADD with an 87% reformate yield which 

would be consistent with a higher pressure, semi-regenerative catalytic reformer. There is a reformate yield chart in 

Petroleum Refining1that matches this data for a feed quality with N+2A = 60 (N= Vol.% naphthenes, A= Vol.% aromatics). 

Following this curve shows the reformate yield effect for operating at higher severities. Continuous catalytic reformers (CCR) 

operate at lower pressures, which improve reformate yields. In our modeling work, Stancil generally assumes about 

2.5 Vol.% improvement in reformate yields with a CCR unit. The table below summarizes the approximate yields, as

discussed. 

Semi-Regenerative Catalytic 

Reformer

Continuous Catalytic 

Reformer

Reformer Severity

(RON)

Reformate Yield

(Vol. %)

Reformate Yield

(Vol. %)

90 87.0 89.5

95 84.5 87.0

100 78.0 80.5

To attain the amount of reformate produced, the actual reformer throughput will need to be multiplied by the appropriate yield 

assumption in this table. For refiners making premium grade gasoline, reformers generally need to operate at a 97 RON to 

99 RON severity for some amount of time to make a blending component with high enough octane to meet premium gasoline 

specifications of 91 to 93 road octane number. The rest of the time, they may only need to operate in the range of 92 RON to 

 

1

James H. Gary and Glenn E. Handwerk, Petroleum Refining: Technology and Economics, 4th ed., Marcel Dekker, Inc., New York, 2001, p. 202.
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V. EPA Inquiry B

Stancil & Co.____________________________________________________________________________________________________________________________________________

V-2

95 RON as needed, to blend low octane components such as LSR, isomerate, light hydrocracker gasoline, light reformate, 

and raffinate. 

Refiners in the business of producing aromatics for petrochemical feedstocks may run their reformers at full rates and high 

severities all the time to maximize production of benzene, toluene, and xylene compounds. Also, some refiners that do not 

have on-site hydrogen production facilities or have access to off-site supplies of hydrogen may also run additional reformer 

throughput and severities to generate hydrogen for their desulfurization processes. 

In general, Stancil believes that assuming the reformate yields at 90 RON currently used in the EPA model may be 

overestimating the volume of reformate production. While we have no data to show what the average might be for the entire 

U.S., the refineries we work with on a routine basis typically operate in a range of 94 RON to 98 RON for semi-regenerative 

reformers and up to 100 RON for CCR units depending on their specific octane situation. Most of these refiners operating at 

the high end of the severity spectrum are undercutting reformer feed to distillates or blending reformer feed in gasoline. The 

refiners operating on the low end of the severity range tend to be more limited in their distillate operations and are reforming 

full range naphtha (i.e., 210°F to 380°F).

Based on Stancil's view of reforming operations as described above, we would agree in principal with the concepts presented 

in the “Reformer Feed Logic” section in the MSAT2 representations of the EPA model for balancing reformer feed and overall 

gasoline volumes.

Logically, we believe that shifting the 340°F to 400°F naphtha into jet fuel would be the first step to reducing reformer feed. 

The second step would then be to shift the 180°F to 285°F naphtha into gasoline blending or feedstock sales. Even so, it may 

not be reasonable to assume that refiners have enough excess octane in their pool to be able to blend large volumes of low 

octane 180°F to 285°F naphtha. We would recommend that EPA develop an octane balance for the gasoline pool to validate 

the assumptions for blending low octane reformer feedstock into gasoline. These octane balances would also shed some 

light on the appropriate reformer severity to assume for each refinery.

In practice, refiners typically do not have a separate distillation tower to separate the reformer feed into 180°F to 285°F and 

285°F to 340°F naphtha fractions. For the 180°F to 285°F naphtha fraction to be blended into gasoline or sold, it would need 

to be fractionated overhead with the LSR. For those refineries operating an LSR isomerization unit, cutting the LSR with a 

285°F cut point could be problematic from a catalyst cycle length standpoint. Also, the cut point temperature between 

reformer feed and kerosene can range from as low as 300°F when operating in maximum kerosene mode to a high of 380°F 

when operating in maximum gasoline mode. Stancil has no data to indicate what the average HSR naphtha cut point is at this 
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Stancil & Co.____________________________________________________________________________________________________________________________________________

V-3

point for U.S. refiners, but would offer that the refineries we work with that are undercutting reformer feed to jet fuel typically 

operate with a HSR cut point in a range of 320°F to 350°F.

Adjustments to reformate volume are needed to account for those refineries extracting aromatics for petrochemical feedstock 

sales or selling the aromatic extraction unit feedstocks. Reformate is usually split into a light aromatics-rich cut called "light 

reformate" that contains the benzene, toluene, and xylene aromatic compounds and is the feedstock for the aromatics 

extraction process and heavy reformate. Light reformate may contain 60 Vol.% to 70 Vol.% aromatics. The remaining heavy 

reformate is a high octane component that is sent to gasoline blending. After extraction of the aromatics from the light 

reformate, the remaining non-aromatic material called "raffinate" is routed to gasoline blending. With all of the aromatic 

materials removed, raffinate is a very low octane high vapor pressure blending component. Some refineries sell or transfer 

the light reformate to off-site aromatic extraction plants, which would show up as a sale or intercompany transfer of 

petrochemical feedstock. Usually the raffinate is returned to the refinery for gasoline blending if the extraction plant is nearby. 

In the EPA model, we assume that the total volume of the aromatics plant throughput reported by OAQPS represents the 

volume of total feedstock (light reformate) going to the aromatics extraction unit. The volume of aromatics and aromatic plant 

feedstocks sold appear to be represented in two columns on a worksheet titled “T3” in the model. One labeled as “Aromatics” 

matched the B/SD values in the EIA refinery capacity report for 2011 and the other labeled as “2009 shipments exPetch,” 

which included a comment that the column was no longer in use. The volume for aromatics sales also appear on a worksheet 

titled “Data For 2011” in a column labeled “Aromatics to Petrochemicals - KBPCD”. The volume of aromatics sold in this 

column is appropriately subtracted from the total reformate volume that is used in the gasoline volume balance. However, it 

did not appear that the volumes of aromatics and aromatic feedstock sales shown in worksheet “T3” were transferred to the 

aromatics sales column in the worksheet “Data For 2011” for the refineries Stancil did calculations for.
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VI. EPA Inquiry C

Stancil & Co.____________________________________________________________________________________________________________________________________________

 VI-1

Inquiry C – Comment on EPA incorporating, and how EPA incorporated in its refinery-by-refinery cost model, refiner plans for 

complying with the Mobile Source Air Toxics [MSAT2] rulemaking to reduce the content of benzene in their gasoline. This 

affected the volume of benzene precursors sent to the reformer or the volume of benzene extracted from the gasoline pool.

We noted that the EPA model initially used a cut point temperature of 160°F for splitting LSR and HSR. For MSAT2

compliance, the EPA model later assumes the LSR cut point shifts up to 180°F to reduce the volume benzene and benzene 

precursors going to the reformer.

In practice, we have found that many refiners operate their fractionators to a LSR cut point temperature of 200°F to 210°F for 

benzene control. While benzene and its precursor molecules boil in the range of 155°F to 180°F (shown in the table below), 

many fractionators used for splitting LSR and HSR simply do not have the number of trays and efficiency that would be 

needed to produce a sharp cut at 180°F that will remove most of the benzene and benzene precursors. As a result, there will 

be some benzene and benzene precursors that overlap into the higher boiling range fractions that diminish as the boiling 

range goes higher. Most compounds that boil in the 180°F to 200°F range do not have characteristics that make good 

reforming material, so there is little incentive to spend the capital or extra energy needed for recovery as reformer feed. 

Compound Boiling Point, °F

n-Hexane 155

Methylcyclopentane 161

Benzene 176

Cyclohexane 177

Conversely, refiners in the business of aromatics recovery may operate below a 150°F cut point between LSR and HSR in 

order to maximize benzene production in their reformers.

We noted that the EPA refining model was generating much more reformer feed for the total U.S. than was indicated by the 

actual reformer rates. After various adjustments in the model’s MSAT2 section, it appears that most of the 160°F to 180°F 

boiling range material was removed from reformer feed and routed to gasoline blending, isomerization, or sales. Expanding 

this MSAT2 cut point adjustment range to include the 180°F to 200°F fraction would directionally help reduce this reformer 

feed imbalance problem. The resulting increase in LSR production may also directionally increase the cost of compliance for 

the Tier 3 regulations.
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Stancil & Co.____________________________________________________________________________________________________________________________________________

 VI-2

As mentioned in the previous section, final adjustments to reformate volume in the EPA model need to account for aromatics 

sales as petrochemical feedstocks as well as sales of light reformate that are used for off-site aromatic extraction plant feed.

A benefit for refineries having these operations is that little else is needed to comply with MSAT2 regulations.

For refineries utilizing isomerization or benzene saturation to control finished gasoline benzene levels, the reformate stream is 

also split, but the light reformate cut is only deep enough to capture the benzene. Typically, light reformate in this mode of 

operation has a cut point of 210°F to 220°F that allows the remaining heavy reformate to have benzene levels below 

0.6 Vol.%. This very light reformate stream can be feedstock for the isomerization or benzene saturation units. The toluene 

and xylene materials remain in the heavy reformate fraction that goes on to gasoline blending. Light reformate has a low 

octane rating, generally in the range of 70 road octane number (RdON) to 75 RdON depending on benzene content.

Saturating the benzene will generally lower the octane rating by 3 RdON to 4 RdON.
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VII. EPA Inquiry D

Stancil & Co._____________________________________________________________________________________________________________________________________________

 VII-1

Inquiry D – Review the methodology applied by EPA to estimate that refiners are maximizing propylene production at the 

expense of FCC naphtha production. Using refinery-by-refinery propylene sales information provided by EIA, EPA estimated 

that higher amounts of propylene production compared to the feedstock volume to the FCC unit would have caused lower 

FCC naphtha production. 

Propylene demand in the U.S. is about 15 million metric tons (MT) per year. The traditional sources of propylene have been 

naphtha-fed steam crackers that produce ethylene for the petrochemicals industry and FCC units at oil refineries. Together 

these sources supply about 90% of the propylene market. For both of these processes, propylene has historically been a byproduct resulting from the production of the primary products these units were designed to make. The table below puts the 

demand for these various products into perspective. 

Product

U.S. Demand

(Million MT/Year)

 Propylene 15

 Ethylene 129

 Gasoline 26,600

Due to the low relative demand, propylene supply availability is generally more dependent on the demand for ethylene and 

gasoline than its own market. For example, if gasoline demand is high, refiners may increase run rates on their FCC units to 

produce more gasoline which also produces more propylene. Since propylene can also be used as feedstock for alkylation 

units, refiners might also choose to divert some propylene from sales to producing alkylate for gasoline blending or sales. 

Conversely, if demand for gasoline declines, refiners will cut FCC runs to reduce gasoline and consequently propylene drops 

as well. 

The graph below is a plot showing the average refinery production of propane and propylene by U.S. refiners since 1990. 

This data combines the volumes of propane with refinery-grade propylene (which is a usually an 80/20 mix of 

propylene/propane). As can be seen, refinery production of propane and refinery-grade propylene sales peaked between 

2000 and 2004 and has been in a slight down trend since then, losing roughly 30,000 barrels per day (B/D) to 40,000 B/D on 

average. 
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 VII-2

The next graph is a plot of the average refinery FCC unit charge over the same time period. FCC unit charge rates peaked 

around 2004 and were in decline until 2010 and have since appeared to stabilize. On average the decline has been around 

400,000 B/D. Using a refinery-grade propylene yield of 11 Vol.% on FCC feed would translate into a production loss of 

40,000 B/D to 45,000 B/D refinery-grade propylene. These charts do not indicate that refiners were taking any extraordinary 

measures to maximize propylene production beyond what they had been doing in the past.

350

400

450

500

550

600

650

19901991199219931994199519961997199819992000200120022003200420052006200720082009201020112012

Thousands of Barrels Per Day

EIA - U.S. Refinery Production of Propane and 

Propylene
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 VII-3

The chart below shows refinery-grade propylene and unleaded gasoline prices on the USGC since 2009. As can be seen, 

refinery-grade propylene prices were very strong relative to gasoline from about mid-2009 through the summer of 2011. 

Based on the strength of these prices, we would opine that polypropylene producers may have bid up propylene prices to 

divert extra volume away from alkylation and gasoline production. 

4200

4400

4600

4800

5000

5200

5400

5600

19901991199219931994199519961997199819992000200120022003200420052006200720082009201020112012

Thousands of Barrels Per Day

EIA U.S. Fluid Catalytic Cracking Unit Charge

0

50

100

150

200

250

300

350

400

2009 2010 2011 2012 2013

USGC, Cents Per Gallon

USGC Refinery-Grade Propylene and Unleaded 

Gasoline Prices

Refinery-Grade Propylene Unleaded 87
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Stancil & Co._____________________________________________________________________________________________________________________________________________

 VII-4

The same trend is shown on the chart below which compares the prices of refinery-grade propylene with alkylate on the 

USGC. Since the end of 2011, propylene prices have been very volatile which would indicate supply and demand balance is 

very tight. 

From the standpoint of the methodology used in the EPA refining model, we would agree that refiners were likely doing 

whatever they could to maximize propylene for sales for much of 2011 based on the strength in propylene prices. This would 

include adding ZSM-5 catalyst additive to their FCC units and diverting propylene feed away from alkylation to sales. For 

2012 and 2013, the weaker price structure might have seen an on-again, off-again situation for maximizing propylene which in 

and of itself could have contributed to the price volatility. Refiners assess the economics for these operating decisions on a 

monthly basis and implementing either strategy is relatively easy to start and stop as economics dictate. ZSM-5 catalyst 

additive has been used in this manner for the last 25 years. 

Generally, we see base refinery-grade propylene yields of 10 Vol.% to 12 Vol.% on FCC feed without ZSM-5. Assuming a 

recovery factor of 0.85 would leave 8.5 Vol.% to 10 Vol.% for propylene sales. The EPA model is currently set up to activate 

the FCC naphtha to propylene volume shift at 6.5 Vol.%. We would recommend raising the activation level to at least

8.5 Vol.% if the EPA decides to continue using this shift.

0

50

100

150

200

250

300

350

400

2009 2010 2011 2012 2013

USGC, Cents Per Gallon

USGC Refinery-Grade Propylene and Alkylate 

Prices

Refinery-Grade Propylene Alkylate
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Stancil & Co._____________________________________________________________________________________________________________________________________________

 VII-5

Generally, each refiner has some maximum amount of ZSM-5 they can add to their FCCU that will allow them to capture 

some of the propylene and butylene yield benefits and still stay within the unit’s light ends handling capacities. Being a small 

market with unpredictable pricing margins would deter many refiners to expand this equipment for propylene alone. Those 

refiners situated near a petrochemicals complex that consumes propylene and could lock in long-term contract agreements 

would be more likely to upgrade their facilities to maximize propylene. For most of our modeling work, we have been using up 

to 5 weight percent (Wt.%) ZSM-5 concentration in FCC catalyst. 
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VIII. EPA Inquiry E

Stancil & Co._____________________________________________________________________________________________________________________________________________

 VIII-1

Inquiry E – Comment on EPA’s methodology of forcing each refinery’s gasoline volume modeled by the refinery-by-refinery 

cost model to match actual refinery gasoline production volume as reported by refiners to EPA. In trying to match individual

refinery gasoline volumes, we used the practice of undercutting FCC naphtha and heavy naphtha into the diesel and jet fuel 

pools. Since we often had mismatched gasoline volumes in those refineries with hydrocrackers, we also estimate

hydrocracker operation (naphtha, intermediate, or diesel modes) as a means to match gasoline volumes. More often than 

not, heavy naphtha volumes tend to exceed reformer throughput volumes, so for those refineries that have excessive gasoline 

volumes, we assume that the excessive heavy naphtha volume is sold. For refineries with insufficient gasoline volume, this 

excess heavy naphtha volume is assumed to be blended into gasoline (but not reformed). In a couple of cases where there is 

a large shortfall in feedstock for reformers we assume that the heart cut of the FCC naphtha is being sent to the reformer for 

producing more aromatics for aromatics extraction.

We agree that the technique used in the EPA model of undercutting heavy naphtha into jet fuel is a realistic way to balance 

both the reformer feed volumes as well as actual finished gasoline volumes to the model’s theoretical estimates. As 

mentioned in Section VI, EPA Inquiry C, we believe that light gasoline volumes (LSR, hydrocracker (HDC) light gasoline, and 

light coker gasoline) are likely understated with the cut points being assumed and are contributing some of the excess heavy 

naphtha reformer feed being predicted in the model. Undercutting heavy naphtha to the jet fuel or diesel pools is one of the

easiest and most common steps refiners can take to shift gasoline to distillates. 

We also concur that the technique used in the EPA model of undercutting FCC naphtha into the diesel pool is another realistic

way to balance gasoline volumes to the model’s theoretical estimates. In practice, refiners accomplish this by changing the 

cut point between naphtha and light cycle oil (LCO) at the FCC unit’s main fractionator. LCO is usually used as feedstock for 

the diesel hydrotreater (DHT) or HDC units. If a refiner is producing residual fuel oil, LCO may also be used as a blending 

stock to adjust density and viscosity specifications. 

While we agree that refiners have been directionally shifting HDC operations away from naphtha production to produce more 

distillates, it is difficult to generally categorize how much of a shift between naphtha and distillate mode a particular refinery is 

capable of. Determining the yields of naphtha or distillate a HDC will make depends largely on the type of feedstock, 

operating pressure, and catalyst type the unit is operating with. HDC feedstocks can be generally categorized as distillate, 

gas oil, and residual. 

Many older HDC units were originally designed to operate with distillate feedstocks in a boiling range of 600°F to 800°F such 

as diesel, atmospheric gas oil, light vacuum gas oil, and FCC LCO. LCO is typically limited to 25 Vol.% or 30 Vol.% of the 

total feed due to the amount of reaction heat it releases in an HDC. The hydrogenation reactions that occur in an HDC unit 

are exothermic, that is they release heat into the system, so maintaining control of the reactor temperature is a factor in 
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VIII. EPA Inquiry E

Stancil & Co._____________________________________________________________________________________________________________________________________________

 VIII-2

considering feedstocks. Operating pressures on HDC units can range anywhere from 500 pounds per square inch gauge 

(psig) to 3,000 psig. Higher pressure will promote the hydrocracking reaction. The annual O&GJ World Refining Survey

designates HDC units operating over 1,450 psig as being high pressure units. Many of the distillate HDC units listed in the 

O&GJ are considered to be high pressure conventional units. In a full conversion (naphtha) operating mode, a distillate 

feedstock HDC may yield 100% or more of gasoline boiling range material depending on the density of material used as 

feedstock. HDC units able to process heavy VGO as feedstock may be able to yield up to 108% gasoline range material in a 

full naphtha mode operation. With LPG recovery, HDC liquid yields can often be 120% to 130% of feed that result from 

cracking heavy materials into lighter materials and saturating them with hydrogen. 

HDC units operating in a maximum distillate mode usually operate with a different catalyst than those operating in a naphtha 

mode. Fractionating towers on many of the older HDC units were not designed with jet fuel or diesel draw trays or rundown 

lines and require installation of this equipment to operate in a distillate mode. In a full distillate operating mode, a distillate 

feedstock HDC may only yield 10% or less of gasoline boiling range material with distillate yields of 95% or more depending 

on the density of material used as feedstock. HDC units able to process heavy VGO as feedstock may yield 15% gasoline 

range material in full distillate mode operation. With LPG recovery, HDC liquid yields may only be 105% to 110% of feed. 

Losing 15% to 20% of liquid recovery volume can be a large disincentive for a refinery to operate in a full distillate mode. 

Refiners we are familiar with tend to operate in some intermediate mode between the two cases described. 

Some HDC units with heavy VGO feed operate at mild to moderate pressures (below 1,450 psig) and only partially convert 

the feedstock to gasoline and diesel products. The remaining unconverted material is processed by the FCC unit. This 

process is similar to an FCC feed treater. The difference is that instead of hydrotreating catalyst, hydrocracking catalyst is 

used. Some existing FCC feed hydrotreaters have been converted to this mode of operation. Yields of gasoline range boiling 

material for these units will be in the 10% to 15% range, distillate yields range from 30% to 35%, and 65% to 55% of 

hydrotreated FCC feed depending on the conversion. 

In summary, a lot of effort is spent in the EPA model to balance theoretical yields of gasoline against actual production 

volumes by shifting gasoline into distillates. Some of these are undercutting reformer feed into jet fuel, undercutting FCC 

naphtha into light cycle oil (LCO), and shifting hydrocracker (HDC) operations from naphtha mode into diesel mode. All of 

these are legitimate means that refiners use to shift gasoline production to jet fuel or diesel production. In the model, these 

shifts are done as needed to force the gasoline and reformer feed to balance with actual throughput and production values for 

each refinery though there is no data available to verify that a particular refiner is actually operating the way that the model 

assumes. One way the EPA could lend credence to the methodology would be to perform a distillate balance to see if the 

theoretical yields match actual data. 
Environmental Protection Agency

Peer Review of Refining Industry Cost Model

IX. EPA Inquiry F

Stancil & Co.____________________________________________________________________________________________________________________________________________

 IX-1

Inquiry F – Review the new data received from vendors and how EPA is using that data. As suggested by the first peer 

reviewers, we requested and obtained more information from vendors of gasoline desulfurization equipment and included this 

information in the final rule cost analysis. The vendors confirmed that the hydrogen consumption values that they reported 

were actual, not stoichiometric.

Vendor 1

Vendor 1 supplied inside battery limits (ISBL) capital and operating cost data for two different sets of cases where the initial 

FCC naphtha sulfur concentrations were assumed to be 2,000 ppm, 800 ppm, and 200 ppm. The basis for each of these 

cases was at 30,000 B/D and assumes each of these naphtha streams are treated down to 75 ppm sulfur with a single-stage 

reactor to represent the current Tier 2 FCC naphtha sulfur levels. The investment and operating cost data supplied for each 

case represents the incremental costs to further treat the FCC naphtha from 75 ppm to 25 ppm sulfur or from 75 ppm to 10 

ppm sulfur. 

The first set of cases provided a wide array of minimum investment options for treating FCC naphtha from 75 ppm sulfur down 

to 25 ppm and 10 ppm by applying some undisclosed minor facility modifications and operating the existing single-stage 

reactor at more severe conditions to meet the 25 ppm and 10 ppm sulfur specifications. Most of these cases result in higher

octane losses and shorter run lengths between catalyst change-outs. The vendor also indicated that in addition to additional 

catalyst costs and the associated maintenance, there would also be costs associated with storing untreated gasoline or 

shutting down the FCC unit. We would also add that shutting down or reducing rates significantly on the FCC would also be 

expected to curtail rates on virtually every unit in the refinery starting with the crude unit which produces most of the FCC feed 

and coker unit which receives its feed from the crude unit. 

Ideally, ancillary units like FCC naphtha hydrotreaters should be designed so the catalyst run length coincides with the normal 

turnaround schedules of the primary operating units (crude, FCC, coker, HDC, etc.) they serve. For an FCC unit, the 

turnaround cycle is usually four to five years. While we understand there may be some potentially new developments in 

technology and catalyst design that may improve the performance of existing FCC naphtha hydrotreaters, there was no 

information available in Vendor 1’s information package to evaluate the basis of these claims. Therefore, while these 

minimum investment cases are interesting, we are not convinced that they are practical from an operations standpoint. 

We noted the EPA incorporated some of this data in the refining cost model for evaluating FCC naphtha treating costs 

associated with the 10 ppm gasoline standard, but we were unable to follow how the vendor data might have been interpreted 

to get the investment requirements shown for the 200 ppm and 800 ppm cases. For the 200 ppm case, the EPA model shows 

$157.5 per B/D investment versus the vendor data that indicates a $0 investment for a 75 ppm to 25 ppm FCC naphtha sulfur 
Environmental Protection Agency

Peer Review of Refining Industry Cost Model

IX. EPA Inquiry F

Stancil & Co.____________________________________________________________________________________________________________________________________________

 IX-2

reduction. Operating costs used in the model for this case matched the data supplied by the vendor. For the 800 ppm case, 

the EPA model shows $60 per B/D investment versus a range of $0 to $100 per B/D investment for a 75 ppm to 25 ppm FCC 

naphtha sulfur reduction. Operating costs used in the model for this case matched the data supplied by the vendor. 

We note that one of the first peer reviewers had similar concerns with the apparent inconsistency of investment requirements 

between the 200, 800, and 2,000 ppm cases in their 2011 report. If the EPA plans to use this minimum investment data, we 

would recommend that the additional costs associated with the FCC throughput reductions or shutdowns recommended by 

the vendor be included in the EPA model analysis. 

The second set of ISBL and operating cost data provided by Vendor 1 assumes installation of a second-stage reactor in the 

800 ppm and 2,000 ppm sulfur FCC naphtha feed cases and a gasoline splitter on the 200 ppm sulfur FCC naphtha feed 

case. These cases applied primarily to reducing FCC naphtha sulfur to 10 ppm for the 5 ppm sulfur gasoline standard. All of 

this data for these cases was incorporated into the EPA model. Vendor 1 did include one case that added a second-stage

reactor on an 800 ppm sulfur FCC naphtha feed case that reduced FCC naphtha from 75 ppm to 25 ppm for the 10 ppm 

gasoline standard but was not used in the EPA model. This seemed like a more reasonable case to use than the one chosen. 

Vendor 1 also provided ISBL capital and operating costs for installing a grassroots post-treater to reduce FCC naphtha from 

100 ppm to 10 ppm sulfur to apply in situations where there is no existing FCC gasoline treater. We noted the EPA model 

also used this same capital cost data for the case to reduce FCC naphtha from 100 ppm to 25 ppm sulfur. The difference 

between the two cases was that the operating costs for the 25 ppm case were lowered proportionally from the 

10 ppm case to reflect less sulfur removal. 

We noted that Vendor 1 did not give an accuracy assessment for the quotes given. Scoping type estimates such as these 

typically have +/- 30% to 50% accuracy, as Vendors 2 and 3 stipulated for their estimates. We note that the EPA model is 

currently set up not to include any capital cost contingency. We would recommend that the EPA consider adding at least a 

30% capital cost contingency factor to the model. Vendor 1, however, did indicate outside battery limits (OSBL) costs to be 

30% to 50% of ISBL investments versus the EPA model assumption of 35% of ISBL. 
Environmental Protection Agency

Peer Review of Refining Industry Cost Model

IX. EPA Inquiry F

Stancil & Co.____________________________________________________________________________________________________________________________________________

 IX-3

Vendor 2

Vendor 2 supplied ISBL capital and operating cost data to upgrade existing units that are operating to reduce FCC naphtha 

sulfur concentrations of 800 ppm to 75 ppm. Two cost estimates were provided for installing a new polishing reactor to allow 

taking the treated FCC naphtha sulfur level from 75 ppm to 25 ppm and 75 ppm to 10 ppm. The ISBL capital costs in the EPA 

model for these two cases were consistent with the vendor estimates. The octane losses shown in the EPA model were 

somewhat higher than the vendor estimates. Vendor 2 provided estimates for hydrogen, fuel, and power consumption as a 

percentage increase over existing consumption, but did not give values for the existing consumption. Since the EPA model 

has data in per barrel consumption values for these items, we assume the EPA had access to the existing operating cost 

data.

Since Vendor 2 did not supply cost estimates for either the 200 ppm or 2,000 ppm FCC feed sulfur cases, we assume the

EPA applied some method to scale the 800 ppm FCC feed sulfur cost estimate up or down to generate the capital and 

operating cost estimates for the 200 ppm and 2,000 ppm FCC feed sulfur cases shown in the model. We were not able to 

readily discern how these adjustments were performed; however, the costs shown appeared to be directionally what we would 

expect. As mentioned previously, Vendor 2 indicated the accuracy of their capital cost estimates were +/- 50%. We would 

reiterate our recommendation that the EPA use a 30% contingency factor for capital cost estimates in the EPA model.

Vendor 3

Vendor 3 supplied ISBL capital and operating cost information for both single-stage and two-stage FCC naphtha hydrotreaters 

processing FCC naphtha with sulfur levels of 800 ppm at 10,000 B/D. The single-stage cases provided the costs for reducing 

the FCC naphtha sulfur from 800 ppm to a base case level of 75 ppm and also to 10 ppm. The two-stage cases provided the 

costs of reducing the FCC naphtha sulfur from 800 ppm to a base case level of 25 ppm and also to 10 ppm. 

To estimate the costs for adding a second-stage reactor, the EPA subtracted the capital and operating costs for a two-stage 

unit from the costs of the base case single-stage unit (800 ppm to 75 ppm sulfur). This was done for both the 800 ppm to 

25 ppm and the 800 ppm to 10 ppm two-stage reactor cases. The resulting differences in capital cost were multiplied by a 

vendor supplied factor of 1.15 to account for extra costs that might be incurred from adding the second reactor to an existing 

unit versus building a new two-stage unit from scratch. The resulting capital costs and operating costs for each of the new 

second-stage units appear to give reasonable results. We would still recommend that the EPA use the +/- 30% contingency 

factor provided by the vendor for capital cost estimates in the EPA model.
Environmental Protection Agency

Peer Review of Refining Industry Cost Model

IX. EPA Inquiry F

Stancil & Co.____________________________________________________________________________________________________________________________________________

 IX-4

We were able to validate Vendor 3’s cost estimate for building a two-stage unit based on the construction costs we have seen 

working with various clients and construction companies. Including the 30% contingency factor, Vendor 3’s cost estimates for 

building a two-stage FCC naphtha hydrotreater match fairly well with what our own cost curves would predict. While we do 

not have any cost curves for adding a second-stage reactor alone, the fact that we can match the vendor’s capital cost for a 

full two-stage unit allows us to have confidence in the method used by the EPA for estimating Vendor 3’s costs for the 

second-stage reactor.
Environmental Protection Agency

Peer Review of Refining Industry Cost Model

X. EPA Inquiry G

Stancil & Co.____________________________________________________________________________________________________________________________________________

X-1

Inquiry G – Review the methodology EPA used to adjust desulfurization costs to account for cases when a refinery’s modeled 

desulfurization situation differed from the typical case for which the vendors provided us information. For example, for 

reducing a refinery’s gasoline sulfur from 30 ppm to 10 ppm, the refiner would typically need to reduce its FCC naphtha from 

75 to 25 ppm. Depending on the amount of FCC naphtha blended into its gasoline, the amount of sulfur control that the 

refiner would need to achieve in its FCC naphtha would be larger or smaller than this. We linearly adjusted the desulfurization 

cost to account for the variances from a typical case.

For cases where the total sulfur removed between the Tier 2 and the Tier 3 case was greater or lower, a factor was generated 

to account for the difference in sulfur removed. For example, if a refinery has an existing FCC naphtha post-treater and the 

total amount of FCC naphtha sulfur removed needed to be 55 ppm to meet the 10 ppm gasoline pool sulfur standard, but the 

amount of sulfur removed by the new treating equipment installed was based on 50 ppm (75 ppm to 25 ppm), a factor of 1.10 

was calculated (55/50 = 1.10) and used to increase the amount of H2 consumption and increase the amount of octane loss to 

account for the additional costs to remove the extra sulfur. Using the same formula, if the amount of sulfur that needed to be 

removed was lower than 50 ppm, the factor would reduce the amount of hydrogen consumption and octane loss.

We would concur that adjusting the H2 consumption and octane loss costs are appropriate. We would also agree that there is 

no need to adjust the fuel, power, and catalyst costs as these are small in comparison. Also, there is no need to adjust the 

capital as this cost is more sensitive to naphtha charge rate than to the amount of sulfur removal. 
Environmental Protection Agency

Peer Review of Refining Industry Cost Model

XI. EPA Inquiry H

Stancil & Co.____________________________________________________________________________________________________________________________________________

XI-1

Inquiry H – Review the methodology EPA used to adjust desulfurization costs to account for situations where the level of 

desulfurization increases above a certain point that causes the desulfurization cost to be increase[d] substantially in a 

nonlinear manner, thus the costs begin to increase exponentially. If we did not make this adjustment, we believe that we 

would be underestimating the cost for those refineries which must achieve a very high percentage of desulfurization.

As it turned out, all five of the refineries we tracked calculations for in the model had FCC naphtha post-treaters and fit the 

example described in Section G. Stancil did note that the formula used to calculate the single-stage correction factor for 

desulfurization amounts higher than 96% was unusual as it multiplies fraction of sulfur removal above 96% by 3. We assume 

that this is the adjustment being made to correct the situation described in your inquiry. 

From what we could gather from the various default options in the model, it appears that for refineries with no existing FCC 

naphtha post-treating, the option is to use the capital cost and operating data for Vendor 1’s grassroots post-treater that was 

designed to reduce FCC naphtha from 100 ppm to 10 ppm sulfur. Since this hydrotreater is operating at level not much 

different than the second-stage polishing units estimated, the operating costs are fairly low versus what would be expected for 

a unit removing over 96% of the FCC naphtha sulfur. If over 96% of the sulfur is being removed to get to a 25 ppm target, the 

FCC naphtha feed would be over 600 ppm to 800 ppm sulfur. With this level of feed sulfur, the operating costs would look 

more like the grassroots single-stage hydrotreater estimates provided by Vendor 3 which are roughly three times higher than 

the operating costs for Vendor 1’s grassroots post-treater. If Vendor 3’s post-treater capital and operating cost data were 

used as the default option, then perhaps the single-stage correction factor above would not be needed. Otherwise, we would 

concur that multiplying the operating costs of Vendor 1’s 100 ppm to 10 ppm by 3 is an appropriate correction.
Environmental Protection Agency

Peer Review of Refining Industry Cost Model

XII. EPA Inquiry I

Stancil & Co.____________________________________________________________________________________________________________________________________________

XII-1

Inquiry I – Review and comment on the conclusions that EPA reached through a conversation with technical experts that 

extractive treating of butane is widely practiced today by refiners and that the sulfur level of butane is under 5 ppm. Thus, no 

additional desulfurization needs to occur for butane. Also, between a review of crude oil assays and the follow-up discussion

with technical experts, it was concluded that extractive treating of light straight run naphtha (LSR) from sweeter crude oils will 

yield a low sulfur level in that stream that would not require additional desulfurization under Tier 3. However, even after

extractive treating of LSR from more sour crude oils, LSR could still contain greater than 10 ppm sulfur that refiners may find 

too high under Tier 3 (this assumes that the LSR is being blended straight to gasoline instead of being hydrotreated before 

being isomerized). 

Stancil would agree that extractive treating of butane is widely practiced today but not necessarily for the reason of lowering 

total sulfur in the product. Butane and other LPG products have copper strip corrosion specifications. Hydrogen sulfide (H2S) 

and other acidic compounds are the primary culprits that cause a failure of this test. Even 1 ppm H2S can result in a failed 

corrosion test. Refiners usually treat butane and other LPG products with caustic or amine compounds to extract the H2S and 

neutralize any acidic compounds. 

There are other sulfur compounds in butane called mercaptans that contribute to the total sulfur content of the stream, but do 

not affect the corrosion test. Since caustic treating will also remove mercaptans from butane, a consequence of meeting 

corrosion test specifications is that the butane will usually end up having lower total sulfur. Testing for free H2S at the 1 ppm 

level is difficult to do, so refiners sometimes measure the mercaptan level of the butane going and coming out of the caustic 

treater to assure some level of mercaptan reduction is taking place. Since H2S will be removed before the mercaptans, 

observing a reduction in the mercaptan levels across the treater will assure all of the H2S is removed. This does not, however, 

necessarily assure that an ultra low sulfur butane product is being produced. We have noted sulfur amounts in blending 

butane that range anywhere from 0 ppm to 30 ppm at the refineries we work with.

We would also point out that specifications for storage of refinery-grade butane at Mount Belvieu, the largest underground 

LPG storage facility on the USGC, allows up to 140 ppm total sulfur. This is also the specification for physical butane that is 

used for settling futures contracts at the Chicago Mercantile Exchange (CME). Some LPG pipeline companies also have this 

same 140 ppm total sulfur while others have a 30 ppm total sulfur specification. In addition to butane that is produced at 

refineries, butane produced at NGL fractionation facilities is also stored at Mont Belvieu and may not be treated to the same

sulfur level as at a refinery. 

Many refiners store excess butane production off-site during the summer months when gasoline RVP specifications are low 

and bring them back to the refinery for gasoline blending in the winter months when RVP specifications are higher. If they are 

storing the butane in a fungible storage facility like Mount Belvieu, the sulfur level of the butane they bring back from storage
Environmental Protection Agency

Peer Review of Refining Industry Cost Model

XII. EPA Inquiry I

Stancil & Co.____________________________________________________________________________________________________________________________________________

XII-2

in the winter could easily be higher than the butane sent to storage in the summer given the 140 ppm sulfur specification. We 

have no data on what the prevailing sulfur levels are in Mont Belvieu butane or other underground storage facilities, but to 

assure that ultra low sulfur butane will be available for winter gasoline blending, it seems to us that refiners may need to incur 

extra expenses to set up dedicated butane storage facilities or potentially add more caustic treating capacity to treat the 

butane being brought back from storage to suitable sulfur levels for Tier 3 gasoline blending. 

Sulfur content in light straight run (LSR) naphtha is primarily mercaptans that come in with the crude oil. Based on what

Stancil observed from the crude assays in our database, sulfur or mercaptan levels in LSR do not seem to be related to the 

amount of total sulfur in crude oil. The graph below is a plot of LSR sulfur versus crude sulfur from the crude assay database 

we have been using. The presence of mercaptan sulfur compounds in LSR seems to be more of a unique feature of 

individual crude oil production fields rather than crude gravity or sulfur content. Stancil has observed this same phenomenon 

when screening crude oils for mercaptan content in jet fuel fractions. 

 

0

100

200

300

400

500

600

0.0 1.0 2.0 3.0 4.0 5.0 6.0

LSR Sulfur, Parts Per Million by Weight

Crude Sulfur, Wt.%

LSR Sulfur Versus Crude Sulfur
Environmental Protection Agency

Peer Review of Refining Industry Cost Model

XII. EPA Inquiry I

Stancil & Co.____________________________________________________________________________________________________________________________________________

XII-3

Based on this, we would have to conclude that there is not a particular crude sulfur percentage cut off point where extractive 

caustic treatment of LSR would stop and hydrotreating would begin. The cut-off point in the EPA model is currently set at a 

crude sulfur content of 1.0 Wt.%.

One possible issue with extractive caustic treatment of gasoline boiling range material is that sulfur removal efficiency is lower 

than for lighter LPG feedstocks. We have only anecdotal evidence from a competing vendor that this efficiency may be as low 

as 90%. At this efficiency level, LSR feed with sulfur concentrations over 10 ppm could become problematic if a refinery were 

depending on a treated LSR sulfur content of 1 ppm for gasoline blending as is in the EPA model. We would recommend that 

the EPA try to verify from vendors what the sulfur removal efficiency is for extractive caustic treatment of LSR materials.

For those refineries that have coker units, the naphtha produced is usually processed with the LSR and HSR such that the 

light coker naphtha will end up combined with the LSR. Even with a light sweet crude slate, light coker naphtha can easily 

have 2,000 ppm to 3,000 ppm sulfur content. Considering the efficiency hurdle needed for extractive caustic treatment to be 

effective, we would recommend those refineries with coker units automatically default to installing new LSR hydrotreating 

capacity if needed. 
Environmental Protection Agency

Peer Review of Refining Industry Cost Model

XIII. EPA Inquiry J

Stancil & Co.____________________________________________________________________________________________________________________________________________

XIII-1

Inquiry J – Ensure the integrity of the new calculations (added since the first peer review) in the refinery-by-refinery cost model 

by working through those equations in the spreadsheet. Check the new equations with sufficient frequency (i.e., one refinery 

in each PADD) to ensure that the refinery model formulas refer to the appropriate cells. Report any errors. 

We tracked the calculations for five different refineries, one per PADD as requested, beginning with unit capacities to end with 

the final cost results. All of the equations appeared to calculate correctly and refer to the appropriate cells as far as we could 

tell. Chasing down the logic loops in the model through redundant capacity tables tended to made this task a very time 

consuming process. The refineries we chose had similar sulfur reduction strategies so we were not able to fully check out the 

calculations for some of the more unusual refinery situations, but did our best to understand the logic employed. 

We noted that one of the refineries we tracked did not show any alkylation rate in the actual throughput data supplied by the 

OAQPS even though that particular refinery has an alkylation unit. Perhaps, the data was not supplied or did not get entered 

at OAQPS, but it may point to a problem with substituting actual data for standard capacity data into the model. We would 

suggest the EPA compare this actual data with the capacity data listed in the EIA and O&GJ capacity reports for consistency.