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Chapter 5 Fuel Program Costs

5.1 Methodology

This chapter provides a summary of the methodology used and the results obtained from 

our cost analyses of the proposed gasoline sulfur control. We start by summarizing the refinery

models used for our analysis. We then describe our detailed methodology for estimating the 

sulfur control costs for our proposed sulfur program followed by the results. We present the 

results from our energy and supply analyses for our proposed gasoline sulfur program. Finally, 

we discuss and compare the results of several cost analyses for various sulfur programs. 

5.1.1 Overview

When we began our planning for estimating the cost of additional reductions in gasoline 

sulfur, we considered two different options. One option for estimating the costs would be to 

utilize a linear programming (LP) model, while the second option would be to develop a 

refinery-by-refinery cost model. While the LP refinery models are necessary and appropriate for 

many analyses, they also have several important limitations of relevance here. When used to 

model the cost of nationwide fuel control programs on the entire refining industry, LP models are 

usually used to model groups of refineries in geographic regions called Petroleum 

Administration for Defense Districts (PADDs). The LP refinery model averages the costs over 

the refineries represented in the PADDs; however, the technology chosen by the refinery model 

would normally be the lowest cost technology found by the refinery model. This may represent 

an unreasonable choice of technologies for individual refineries because of how refineries are 

configured and based on the sulfur control technologies installed for compliance with the Tier 2 

gasoline sulfur program. While the choice of technologies can be limited based on an 

approximate analysis of what mix of technologies would best suit the group of refineries 

modeled in each PADD, this would only provide an approximate estimate of the cost incurred. 

Based on the quality of input data to these LP models and the assumptions made for complying 

with a regulatory requirement, LP refinery models may overestimate or underestimate the 

program costs. For example, an LP refinery model would not be a sensible tool for estimating 

the credit averaging and trading between refineries. This could be partially overcome by 

iterating between PADD refinery model runs, thus estimating the number of credits traded 

between PADDs and estimating the level of sulfur control in each PADD. However, the need to 

make multiple runs per PADD for each case, coupled with the need to run multiple control cases 

for different sulfur standards, would be very time consuming, costly and still would only result in 

approximate estimates of the sulfur levels achieved and the cost incurred.

For this reason, EPA developed a refinery-by-refinery cost model which models the 

capability for each refinery to revamp existing or install new sulfur control technologies 

available to them to reduce their gasoline sulfur levels. Rather than start from scratch, we started 

from a refinery-by-refinery cost model developed by APT (Mathpro) for EPA to estimate the 

cost of benzene control under MSAT2. However, instead of using the representations of benzene 

control technology contained in the model, we obtained information about gasoline 

desulfurization and represented the cost of this desulfurization in the refinery-by-refinery cost 

model. 
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We believe the refinery-by-refinery cost model best estimates the cost of individual 

refineries, especially when considering an averaging, banking and trading (ABT) program and

therefore is the best analysis tool for estimating nationwide costs. However, the refinery-byrefinery cost model cannot estimate certain inputs necessary for estimating costs. Because the 

refinery-specific information is not publicly available, it was necessary to find another way to 

estimate this information. The inputs and outputs from LP refinery cost modeling provide this 

needed information and it was utilized in the refinery-by-refinery cost model. The information 

from LP refinery modeling used in the refinery-by-refinery cost model is described in Section 

5.1.3.

Since the refinery-by-refinery cost model contains confidential business information for 

each refinery, we could not publish the model or present some of the details of the model here. 

Therefore, to ensure its viability the refinery-by-refinery cost model was subjected to peer review 

by two refinery industry consultants. Our review of most of the suggested changes 

recommended by the peer reviewers suggested that there would be little to no change in our 

desulfurization cost estimate (some of the changes would increase the estimated costs, while 

others would reduce the estimated costs). Also, we anticipate making other improvements to the 

cost analysis conducted for the final rule, which would necessitate a second round of peer 

review. Therefore, the peer review comments will be addressed prior to undertaking the cost 

analysis for the final rulemaking along with the other changes that we will be making to our cost 

analysis. The peer review comments are contained in two reports submitted to the docket. 

 The refinery-by-refinery cost model focuses on reducing sulfur from the FCC naphtha 

because of its high sulfur content. To comply with the 30-ppm Tier 2 sulfur control program, 

most refiners installed FCC naphtha hydrotreaters (referred to as FCC postreaters) or FCC feed 

hydrotreaters (referred to as FCC pretreaters) to reduce that unit’s sulfur contribution to their 

gasoline pool. If refiners installed an FCC postreater under Tier 2, we modeled refiners 

revamping those units. However, if refiners relied on FCC pretreaters to comply with Tier 2, we 

assumed that grassroots FCC postreaters would have to be installed at those refineries to reduce 

its gasoline pool down to 10 ppm. However, since adding grassroots FCC postreaters is 

relatively expensive for the amount of sulfur reduction obtained, the ABT analysis we conducted 

avoided many of these types of investments. Refineries with both pre and postreaters today

could achieve further gasoline sulfur reductions less than 10 ppm at a relatively low incremental 

cost and sell the credits to those refiners who are operating refineries which would otherwise be 

faced with grassroots postreater investments. In addition to addressing the sulfur in the FCC 

naphtha, we believe that some refineries may need to reduce the sulfur in light straight run (LSR) 

naphtha. Some refineries might also need to reduce sulfur in butane, although we don’t expect 

refiners to need to address butane sulfur unless they are pursuing a very stringent gasoline sulfur 

standard, e.g., 5 ppm.

To better understand the desulfurization costs, we evaluated several different scenarios or 

cases. For a 10-ppm average sulfur standard, we assessed the costs based on each refinery

achieving the 10-ppm standard with no averaging among refineries, an averaging program which 

assumed intra-company transfers of sulfur credits, and an averaging program which assumed 

nationwide transfers of sulfur credits. To provide credits for averaging and trading under the 10-

ppm average standard, we also evaluated refiners reducing their gasoline sulfur down to 5 ppm. 

Since we had estimated costs for each refinery to get to 5 ppm sulfur, we also report out the cost 
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for a 5 ppm average gasoline sulfur standard assuming no averaging between refineries. The 

costs for the proposed sulfur program are based on a 10-ppm sulfur standard with intra-company

credit transfers. These different scenarios are summarized in Table 5-1.

Table 5-1 Sulfur Control Cases Evaluated for the Proposal

10-ppm Standard 5-ppm Standard

No ABT Program No ABT Program

ABT Program with IntraCompany Credit Transfers

(Proposed Rule Costs)

N/A

ABT Program with 

Nationwide Credit Transfers N/A

5.1.2 LP Refinery Modeling Methodology and Results

Although we used the refinery-by-refinery cost model to estimate gasoline 

desulfurization costs, certain input information was needed to estimate the costs with refineryby-refinery cost model, and without access to detailed refinery-specific information, we relied on 

outputs from our LP refinery modeling. Perhaps the most important input is the cost for making 

up the octane loss that occurs with desulfurization. Certain refinery operations information from 

the LP refinery model was used for estimating the volume of gasoline produced in the refineryby-refinery model, including the utilization factors of individual refinery units, and the 

percentage that straight run naphtha, FCC naphtha and hydrocrackate comprises of the feed 

volume of their respective units. 

 LP refinery models are detailed mathematical representations of refineries. They are 

used by individual refining companies to project how best to operate their refineries. They are 

also used by government agencies, such as EPA and DOE, as well as by refining industry

associations and individual companies, to estimate the cost and supply impacts of fuel quality 

changes. LP refinery models have been used for these purposes for decades and a certain 

protocol has been established to conduct these studies. 

Two different sets of refinery modeling runs from two different LP refinery models were 

used as inputs into the refinery-by-refinery cost model. The refinery-by-refinery cost model 

already contained the utilization factors and gasoline production volumes for individual refinery

units from the analysis conducted by Mathpro for the MSAT2, and we continued to use that 

information for this cost analysis. The gasoline demand is expected to be fairly flat in the future, 

so using the previous refinery modeling work for these inputs will likely have little impact on the 

cost estimate. We plan on updating these inputs for the final rule to reflect more recent refinery

modeling work. 

Additional refinery modeling was conducted using the Haverly GRTMPS refinery model. 

The primary reason for conducting new LP refinery modeling analysis was to estimate the cost 

of making up the octane loss associated with desulfurization as well as estimate how gasoline 

qualities would be affected by the octane recovery to feed into the emissions inventory impact 

analysis discussed in Chapter 7. While the gasoline demand and production volumes are not 
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expected to change in the future, the cost of octane is expected to decrease dramatically due to 

expected much larger use of ethanol under the RFS2 rulemaking. 

The first step in conducting an LP refinery modeling analysis was the development of a 

base case. The base case is a refinery modeling case that calibrates the refinery model based on 

actual refinery unit capacity and input and output data. The base year for this study was the year 

2000 for the Mathpro model and the year 2004 for the Haverly model. Because much of the 

information available for establishing the base case is only available for PADDs of refineries, the 

LP refinery modeling was conducted on a PADD-wide basis. Refinery capacity information

from the Oil and Gas Journal was aggregated by PADD and entered into the LP refinery model. 

The feedstock volumes, including crude oil and gasoline blendstocks, were obtained from the 

Energy Information Administration (EIA) and entered into each PADD’s model. Similarly, 

product volumes such as gasoline, jet fuel, and diesel fuel were obtained from EIA and entered 

into the cost model. The environmental and ASTM fuel quality constraints in effect in the base 

year were imposed on the products. This includes the Reformulated Gasoline program and the 

500-ppm highway diesel fuel sulfur standard, and for the Haverly LP refinery modeling, the first 

year of the Tier 2 gasoline sulfur standard. This information was input into the LP refinery cost 

model for each PADD and each PADD model was run to model the U.S. refinery industry for the 

base year. The gasoline quality for each PADD refinery model was then compared to the actual 

gasoline quality for conventional and reformulated gasoline which is available from the RFG 

database. Each model was calibrated to closely approximate the gasoline quality of each PADD. 

The second step in modeling is the development of a reference case. The purpose of the 

reference case is to model the refining industry operations and cost in a future year, which is the 

year that the control program is modeled to be in effect (serving as a point of reference to the 

control cases for estimating costs and other impacts). The reference year for the Mathpro LP 

refinery modeling was 2012 while the reference year for the Haverly refinery modeling was 2017

and 2030. We developed two reference cases with the Haverly model to model different control 

case scenarios. The reference case was created by starting with the base cases for each PADD 

and adjusting each base case to model the future year, accounting for the changes between the 

two years. 

Two different types of adjustments were made to the base case refinery models to enable 

modeling the refining industry for the reference case. First, the change in certain inputs such as 

product volumes and energy prices need to be accounted for U.S. refinery gasoline, diesel fuel 

and jet fuel demands are projected year-by-year by EIA in its Annual Energy Outlook (AEO); 

the projections from the AEO for the reference case are used in the refinery modeling analysis. 

The Mathpro LP refinery modeling relied on AEO 2006 while the Haverly LP refinery modeling 

relied on AEO 2011. This growth in demand is used to project refinery production for each 

PADD to meet that increased demand. This projected growth in U.S. refinery production was

entered into the reference case version of the LP refinery model. The utility and crude oil and 

other feedstock prices which are projected by EIA for the future year being modeled were also

entered into the refinery model as well as the estimated product prices. 

The second adjustment made to model the reference cases was the application of fuel 

quality changes. Environmental programs which have been implemented or which will largely 

be implemented by the time that the prospective fuels control program would take effect were 
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modeled in the reference case. These fuel quality changes include limits such as the 30-ppm 

average gasoline sulfur standard, 15-ppm caps on highway and nonroad diesel fuel and the 

MSAT2 benzene control program, in addition to the environmental programs which were already

being modeled in the base case. For the Mathpro refinery modeling, which was conducted 

before the nonroad diesel fuel program and MSAT2 benzene control programs were finalized, 

those fuels control programs were not modeled in the reference case. Also, the implementation 

of EPAct required a large increase in the amount of ethanol to be blended into gasoline to 

comply with the renewable fuels standard (RFS), but not RFS2. In its AEO 2006, EIA projected 

that the volume of ethanol blended into gasoline exceeded the RFS required amounts, resulting 

in 9.6 billion gallons of ethanol blended into gasoline by 2012. Other provisions of EPAct that 

were modeled with both the Mathpro and Haverly models included a de facto ban on MTBE and 

rescinding the RFG oxygenate requirement. The reference case unit throughputs and gasoline 

blendstock volumes were used in the refinery-by-refinery cost model. For the Haverly refinery

modeling work, in addition to the EPAct provisions, the RFS2 renewable fuels volumes were 

modeled for 2017. For the 2017 reference case, 17.8 billon gallons of ethanol were assumed to 

be blended into gasoline, and 3.9 billion gallons of renewable and cellulosic diesel fuel and 

biodiesel were assumed to be blended into diesel fuel for the control case. For gasoline, the 

ethanol volume beyond the E10 blendwall was assumed to be blended as E15. For the 2030 

reference case, we modeled 22.2 billion gallons of corn and cellulosic ethanol, and 8.3 billion 

gallons of renewable diesel and biodiesel. 

The third step in conducting the LP refinery modeling was to run the various control 

cases. The control cases are created by applying a specific fuel control standard to each PADD 

reference case. To single out a specific cost or other impact, the sole difference between the 

control case and the reference case is the parameter change being studied. 

For the Haverly modeling, a control case was run to model the octane loss associated 

with desulfurization using 2017 as the year of analysis. Since we solely wanted to identify the 

cost of recovering lost octane for the refinery-by-refinery modeling, this case was run by 

reducing the octane value of the FCC naphtha by one octane number, and this was the sole 

change relative to the reference case. The control case was run with capital costs evaluated at a 

15 percent rate of return on investment (ROI) after taxes.A The octane cost estimated by the LP

cost model is 0.76 cents per octane number per gallon of FCC naphtha. Because the octane loss 

associated with a specific technology may be lower or higher than 1 octane number, we scaled 

the octane cost based on the relative estimated octane loss on the FCC naphtha (i.e., a 1⁄2 octane 

loss of the FCC naphtha was estimated to cost 0.38 cents per gallon of FCC naphtha. Table 5-35

at the end of this chapter summarizes the data output from the refinery modeling from which we 

calculated the octane cost for using in the refinery-by refinery cost model. 

It was necessary to estimate the gasoline qualities for estimating the emissions impact of 

the proposed Tier 3 program. This was conducted in two separate steps. First it was necessary

 

A

 Normally we conduct the refinery modeling assuming an after-tax 15% ROI and adjust the costs to reflect a 

before-tax 7% ROI to report the costs. However, in this case because the new capital investments were so minimal, 

we omitted the capital cost amortization adjustment because its effect on costs was judged to be negligible. 
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to estimate the gasoline qualities of the 2017 and 2030 reference cases relative to the gasoline 

qualities of a revised base case. The sole differences that we modeled between the 2005 revised 

basecase and the 2017 and 2030 reference cases was the phase out of MTBE and the addition of 

ethanol. For the 2005 revised basecase we modeled 1.7 billion gallons of MTBE and 4.1 billion 

gallons of ethanol. For the 2017 and 2030 reference and control cases, we modeled 17.8 and 

22.2 billion gallons of ethanol, respectively. In 2017, we estimated that approximately half the 

gasoline would be 10 percent ethanol and the about the other half would be 15 percent. To 

model the emissions impact of the different ethanol blends, we modeled two reference cases, one 

with 100 percent E10 and the other with 100 percent E15. These two ethanol cases were 

modeled in 2030 and we used the results for 2017 as well. The gasoline qualities for the 

reference and two ethanol cases are summarized in Table 5-57 to Table 5-61 in the appendix at 

the end of this chapter. The changes in gasoline quality are summarized in Table 5-2. Because 

of the tendency for the LP refinery model to shift gasoline blendstocks around resulting in odd 

gasoline quality changes in individual PADDs, we solely used the national average change in 

gasoline qualities and applied those changes for all E10 or E15 gasoline for the emissions 

analysis.

Table 5-2 Difference in Gasoline Qualities between E10 and E15 Control Cases with the 

Reference Case

Summer Winter Summer Winter

E200 5.08 2.16 11.08 9.71

E300 2.41 -1.89 3.58 3.23

Aromatics -2.88 -1.20 -5.95 -3.99

Olefins -0.92 -1.26 -1.68 -1.75

E200 -3.29 6.96 1.10 11.07

E300 -2.82 2.10 -4.08 2.68

Aromatics -0.17 -2.70 -2.03 -3.22

Olefins -0.31 -1.26 -1.55 -1.46

E200 6.06 5.65 12.03 15.76

E300 3.27 0.76 3.54 8.51

Aromatics -4.93 -5.86 -8.43 -9.89

Olefins -1.06 -2.56 -2.13 -2.03

E200 8.78 5.19 11.29 10.88

E300 0.95 -0.10 1.58 3.81

Aromatics -2.92 -5.01 -4.71 -7.62

Olefins -1.53 -1.06 -1.87 -1.56

E200 3.46 4.65 8.65 11.77

E300 0.88 0.04 1.08 4.50

Aromatics -2.58 -3.19 -5.28 -5.65

Olefins -0.87 -1.59 -1.80 -1.73

E15

US avg 

minus CA

E10 

PADD 1

PADD 2

PADD 3

PADDs 4

& 5OC

The second step for estimating gasoline qualities was to model the impact of 

desulfurization on gasoline qualities. The total impact of desulfurization on gasoline qualities is 
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5-7

comprised of the reduction in gasoline sulfur, the associated reduction in olefins and the impacts 

of recovering the lost octane. The sulfur reduction is fixed by the standard and the olefins 

reduction is a function of the selectivity of the desulfurization technologies. We reviewed the 

information that we had obtained for the gasoline desulfurization technologies and estimated that 

desulfurizing gasoline from 30 ppm to 10 ppm would result in a 1 percent reduction in olefin 

level. Since we estimated the cost of making up lost octane using the LP refinery model, we 

used that case for estimating the impact of octane recovery on gasoline qualities. The gasoline 

qualities for the reference case and the control case which reflects a 1 octane number loss in the 

FCC naphtha pool are summarized in Table 5-47 to Table 5-51 at the end of this chapter. The 

difference in gasoline qualities between the reference and control cases is summarized in Table 

5-3. Because of the tendency for the LP refinery model to shift gasoline blendstocks around 

resulting in odd gasoline quality changes in individual PADDs, we solely used the national 

average change in gasoline qualities and applied those changes for all gasoline for the emissions 

analysis. After we integrated the gasoline desulfurization information into the refinery-byrefinery cost model, we estimated that desulfurizing gasoline from 30 ppm down to 10 ppm 

would result in about a one-half reduction in FCC naphtha octane ((R+M)/2) number. To 

estimate the changes in gasoline quality from a one-half octane number loss in FCC naphtha that 

we estimated, we divided the gasoline quality changes for one octane number in the FCC 

naphtha by a factor of two resulting in half the changes in gasoline quality that we estimated for 

a one octane number change in FCC naphtha. The second set of columns in Table 5-3

summarizes the gasoline quality changes that we estimated for reducing the gasoline sulfur levels 

from 30 to 10 ppm.B 

 

B

Since we completed the LP refinery modeling to estimate the cost for recovering the lost octane and the associated 

changes in gasoline quality, we found that other Tier 3 refinery modeling studies did not show the same increase in 

aromatics and decrease in E300 (see also 7.1.3.2). We then discovered that the LP refinery model that we have 

licensed to use required some improvements in how the refinery model was characterizing both the light-cut and the 

heavy-cut naphtha from the reformer streams to more accurately estimate the E300 and aromatics content of these 

streams. We have subsequently worked with a contractor to make these improvements to the LP refinery model and 

will reassess the changes in gasoline quality for the final rule analysis. Thus, while our modeling results shown in 

Table 5.3 show a meaningful impact on aromatics and E300, we believe there will in fact be little or no change. 

Note that these improvements are not expected to have any impact on the cost estimates made by the refinery model. 
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Table 5-3 Differences in Gasoline Qualities Between the Control and Reference Cases

Summer Winter Summer Winter

E200 -0.20 -0.27 -0.10 -0.14

E300 -2.03 -0.63 -1.01 -0.32

Aromatics 1.14 0.69 0.57 0.34

Olefins 0.60 0.01 0.30 0.01

E200 -0.19 -0.04 -0.09 -0.02

E300 -0.06 -0.12 -0.03 -0.06

Aromatics 0.48 0.27 0.24 0.13

Olefins 0.09 0.03 0.04 0.01

E200 -1.68 -0.06 -0.84 -0.03

E300 -1.99 -0.35 -1.00 -0.17

Aromatics 2.50 0.45 1.25 0.23

Olefins 0.15 0.04 0.07 0.02

E200 0.50 -0.69 0.25 -0.35

E300 0.14 -1.67 0.07 -0.84

Aromatics 0.10 -0.32 0.05 -0.16

Olefins 0.02 -0.49 0.01 -0.25

E200 -1.55 -2.54 -0.78 -1.27

E300 -1.49 -1.37 -0.75 -0.68

Aromatics 1.26 0.95 0.63 0.48

Olefins 0.37 -0.23 0.18 -0.12

US avg 

minus CA

2017 minus 1 ON in Adjusted for 1/2 ON

FCC Naphtha

PADD 1

PADD 2

PADD 3

PADD 4 & 

5OC

5.1.3 Summary of Refinery-by-Refinery Model Methodology

The purpose of the refinery-by-refinery cost model is to project how each refinery would 

reduce the sulfur in its gasoline pool to 10 ppm or lower and to estimate the cost for doing so. 

To do this we created a U.S. refining industry refinery-by-refinery spreadsheet cost model using 

inputs from an LP refinery model case to allow us to better understand the gasoline sulfur control 

costs to individual refineries. This spreadsheet cost model also allowed us to model how costs 

would be affected by an ABT program. 

The building of the refinery-by-refinery model consisted of two major steps. The first 

step was to estimate baseline operating conditions for each refinery. This involves estimating the 

volumes and sulfur levels of the gasoline blendstocks that comprise each refinery’s gasoline. We 

chose to use information from 2009 for modeling the baseline operating conditions for the 

refineries as it’s the latest year we had data for refiner operations and yields. Additionally, EIA 

projections indicated that gasoline demand is expected to be essentially flat between 2009 and 

2017, alleviating the need to adjust refinery operating throughputs and yields for future changes 
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in gasoline demand1.

C

 Because of these factors, the 2009 gasoline production volumes and 

refinery operating conditions can reasonably be projected to be at the same level in 2017 (the 

first year of implementation of the Tier 3 program) in estimating costs and refinery impacts. As 

a final adjustment to our estimated gasoline volumes and sulfur levels, we calibrated the model 

to actual refinery gasoline volume and sulfur levels to ensure our model’s accuracy.

To estimate the cost for each refiner to lower its gasoline pool down to 10 ppm, we used

our refinery-by-refinery model to estimate the FCC naphtha volume, the sulfur level of the FCC 

naphtha, and the amount of sulfur reduction needed in FCC naphtha to meet a 10-ppm sulfur 

standard at each refinery. We also incorporated in our refinery-by-refinery model the impacts 

that FCC pretreaters have on FCC naphtha yields and sulfur levels, as well as the impact of 

refinery-specific crude oil sulfur levels on FCC naphtha yields. Similarly, we also used the 

refinery-by-refinery cost model to estimate the volume levels of light straight run naphtha (LSR) 

and natural gas liquids (NGL) that require additional hydrotreating, as well as butane volumes 

that are directly blended into the gasoline pool.

The second step involves applying the various sulfur control technologies to each refinery

as necessary to meet the 10-ppm sulfur standard. We expect that the majority of the sulfur 

reductions necessary to comply with a 10-ppm gasoline sulfur standard will come from reducing 

the sulfur level in their FCC naphtha. Using our refinery-by-refinery model we also estimate that 

a few refineries will add additional LSR/NGL hydrotreating capacity. We also evaluate each 

refiner’s cost to install new butane Merox extraction equipment to lower the sulfur level of 

butane that is directly blended to the gasoline pool. Reducing the sulfur content of butane was 

assumed necessary to meet a 5-ppm sulfur standard for our ABT cases. This assumption is 

conservative as many refiners may already have this equipment, or may purchase low-sulfur 

butanes that have already been treated by their supplier.

This allows us to generate a cost estimate for the sulfur control technology in each 

refinery. The capital costs for installing the sulfur control technologies in each refinery were 

evaluated based on a 7 percent return on investment (ROI) before taxes. In the following 

sections, we present the various steps that were used in this refinery-by-refinery modeling 

analysis.

5.1.3.1 Estimating Individual Refinery Gasoline Blendstock Volumes

In order to develop a baseline for our refinery-by-refinery analysis, it was necessary to 

understand the sulfur levels and volumes of the various blendstocks which make up each 

refinery’s gasoline. Each refinery blends up its gasoline pool from the various gasoline 

blendstocks that are produced from the refinery units installed at each refinery. However, 

information on the volumes and sulfur levels of each gasoline blendstock produced by each 

refinery is not publicly available, so it was necessary to estimate them. Estimating each 

 

C

Since we conducted the cost analysis for the proposed rulemaking, we put in place an additional round of 

greenhouse emission reductions (2017 – 2025) for light duty cars and trucks that will reduce future gasoline 

demand. When we model the costs for the final rulemaking, we will incorporate this reduction in gasoline demand 

in our costs analyses. 
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5-10

refinery’s gasoline blendstock volumes was accomplished using actual 2009 refinery specific 

throughput rates that we obtained from EIA for crude, FCC, cokers and hydrocracking units, and 

published refinery unit capacity information for the other refinery units. We used this 

information to estimate the extent that each refinery process unit is utilized, followed by a unitspecific analysis for estimating how each refinery unit produces material for blending into 

gasoline. After the unit-by-unit estimates are completed, we performed an overall check by 

comparing our estimated gasoline volumes with reported gasoline volumes for each refinery, 

using EPA’s RFG database

The model requires the total gasoline volume and each gasoline blendstock volume for 

each refinery as an input. Although the model does estimate this volume of gasoline produced 

by each refinery based on the estimated volumes of each gasoline blendstock, we chose to use 

actual 2009 gasoline production data reported by refiners as for the total gasoline volume for 

each refinery in our cost calculations. To comply with the RFG program, refiners report gasoline 

production volumes and sulfur levels for reformulated and conventional gasoline to EPA. We 

used this data and imputed each refiner’s 2009 total gasoline production and corresponding 

sulfur levels into our model.

In the end, our completed refinery-by-refinery modeling estimates of gasoline produced 

on a national basis, correlated very well with the actual refinery production volume in 2009, with 

our estimated volumes having an overall error of approximately 0.5 percent relative to the 

reported refinery production volumes. In order to minimize the impact of this error, we forced 

the estimated total refinery gate gasoline volume to match actual reported 2009 gasoline 

production volume across all the refineries. The unadjusted refinery-by-refinery estimates of 

FCC and LSR gasoline volumes, however, are used for estimating necessary equipment 

modifications and costs for sulfur removal. This is due to the fact that the refinery by refinery

models estimates for each refiner’s FCC gasoline are likely to be very close to actual FCC 

production, as we use actual refinery specific FCC charge throughput rates and account for the 

effects of FCC feed pretreating on FCC gasoline yields. Additionally, the cost for treating FCC 

gasoline in our Tier 3 programs, comprise over 85 percent of the total costs, while LSR 

comprises the bulk of the remaining costs.D 

5.1.3.1.1 Principal Refinery Unit Volumes

To estimate the production volumes for each of the refinery’s gasoline blendstocks, the 

refinery-by-refinery model needs process capacity information. The Oil and Gas Journal (OGJ) 

publishes and the EIA reports unit capacities for the principal refinery units for each refinery in 

the U.S.

2,3

 We updated our database from these two sources to reflect capacity that was in place 

in 2009, the base year for the model. Where differences between the two databases existed, we 

used the information that was judged best overall from the two sources and entered it into the 

refinery-by-refinery cost model. These unit capacities indicate the maximum throughput rate for 

each individual unit, not the actual unit throughput rates for each facility, as this is proprietary 

 

D We did not account for any undercutting of the heavy FCC naphtha into jet and diesel fuel, nor did we account for 

the removal of any pentanes that might be occurring in refineries to comply with stringent summertime RVP 

standards, therefore our analysis is likely somewhat conservative and overestimates the costs. 
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

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business information and not publicly available. In order to enhance our model, we obtained 

from EIA the actual 2009 annual unit throughput rates for each refiner’s crude and major 

refinery units (FCC, cokers and hydrocracker units). With this information, the refinery-byrefinery model was fine-tuned to reflect each refineries gasoline blendstocks yields. Our use of 

this information significantly improved our model’s ability to estimate FCC naphtha, as well as 

other gasoline blendstocks that each refinery makes. The FCC, coker and hydrocracker unit’s 

throughputs versus actual capacity that we obtained from EIA for each domestic refinery on a 

PADD average basis are listed in Table 5-4. This information is presented on a PADD average 

to protect CBI.

Table 5-4 Process Capacity Utilizationa

Crude 

Throughput

FCC 

Throughput

Coker 

Throughput

Hydrocracker 

Throughput

Total U.S. 0.843 0.840 0.761 0.768

PADD 1b0.779 0.754 0.643 0.707

PADD 2 0.859 0.814 0.810 0.774

PADD 3 0.858 0.880 0.782 0.637

PADDs 4/5 

excluding California

0.817 0.794 0.824 0.903

aActual unit throughput rates as a fraction of maximum unit capacity on a PADD basis

b

PADD 1 data includes Hovensa, VI

In the model, we also adjusted the refinery capacity information to account for refinery

expansions or refinery shutdowns that we were aware of and are scheduled to occur over the next 

several years. Refinery expansions include those announced for WRB Refinery in Wood River 

Illinois, the Valero Refinery in Norco Louisiana, and the Marathon Refinery in Garyville, 

Louisiana. For these expansions, there is limited public data on which of the specific process 

unit capacities would be increased, though each expansion project has information on the crude 

unit capacity increase. Since the data was limited, we increased all of the existing individual 

process unit capacities by the fractional increase in crude oil unit capacity at each of the 

expanding refineries. Refiners that we believe are permanently shutdown in PADD 1 were 

removed from our analysis but, consistent with recent import/export trends, we allowed PADD 3 

to supply any lost capacity to PADD 1 as a result of this lost production. PADD 1 refiners that

were presumed to be permanently shutdown are; Giant refinery located in Yorktown, Virginia, 

Sunoco refinery in Westville, New Jersey, and Shell Oil refinery in Bakersfield, California.

5.1.3.1.2 Other Refinery Unit Volumes

The next step was to calculate actual unit throughput rates for the other refinery processes 

that produce gasoline blendstocks. These units include alkylation, dimerization, polymerization, 

isomerization, naphtha reforming. All of these processes feedstocks are primarily supplied by

the crude and FCC Units. Since this data is similarly not publicly available we tuned these units 

to the EIA throughputs rates for crude and FCC units at each facility, with alkylation units 

running at the same throughput rates as the FCC and the remaining units running at the crude oil 

throughput ratesThe results of the capacity utilizations of these downstream units are 

summarized in Table 5-5 below.
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

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Table 5-5 Other Unit Process Capacity Utilizationa

Reformer 

Throughput

Alkylation 

Throughput

Isomerization 

Throughput

Poly/Dimersol 

Throughput

PADD 1 0.774 0.886 0.931 1.000

PADD 2 0.859 0.878 0.859 0.859

PADD 3 0.858 0.880 0.858 0.345

PADDs 4/5 

excluding 

California

0.817 0.794 0.714 0.100

a Actual unit throughput rates as a fraction of maximum unit capacity on a PADD basis

With these inputs the refinery-by-refinery model now contained estimates of the 

feedstock charge rates for all of the gasoline blendstock producing units, though estimating 

refinery unit capacity and capacity utilization may or may not translate directly into the gasoline 

blendstock volume produced by a specific refinery unit. This is because some refinery units may

also produce products other than gasoline blendstock. Additionally, some processes have 

volume loss of feedstock due to process reactions and conversions that take place that increase or 

decrease the density and therefore the volume of products. To take this into account, a gasoline 

fraction yield factor has to be applied to each process to convert the process charge rate into the 

yield of gasoline blendstocks. The process fractional yields that were used in our refinery by

refinery model were taken from our MSAT2 final rule LP refinery modeling work, which 

represented the U.S. refining industry on a PADD basis. The FCC unit process yields of naphtha 

blendstock are different for units with an FCC feed pretreater, versus those without feed 

pretreating. In our modeling we accounted for this by adjusting yields and sulfur levels of FCC 

units with pretreaters and those without a pretreater. The fractional yields of gasoline blendstock 

for the major process units and the 2009 throughputs for each of these units used in our model 

are summarized below in Table 5-6 and Table 5-7.

Table 5-6 Gasoline Blendstock Fraction Yields Per Process Unit Charge

Crude FCC Units 

Average

Coker Hydrocracker

PADD 1 0.190 0.560 0.234 0.369

PADD 2 0.211 0.570 0.234 0.311

PADD 3 0.188 0.554 0.239 0.212

PADDs 4/5 excluding 

California

0.183 0.565 0.234 0.276
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

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Table 5-7 2009 Refinery Unit Throughputs (1,000 BPSD)

Crude FCC Units Coker Hydrocracker

PADD 1 1,624 652 <3 a<3 a

PADD 2 3,193 1,017 322 223

PADD 3 7,262 2,604 1,043 500

PADDs 4/5 excluding 

California

1,363 263 130 94

a

Since there are less than three refiners in this PADD with these units, the data was not reported to protect 

CBI information.

The FCC unit produces significant volumes of naphtha, a gasoline blendstock. The 

conversion percentage to naphtha is affected by the severity of the operation of the FCC unit. As 

shown in Table 5-6 above, the portion of FCC feedstock converted to naphtha ranged from 55 to 

57 percent across the various PADDs. The range among individual refineries can be quite large, 

but we didn’t have access to refinery specific data for this. However, as a group there is 

expected to be differences between refineries with and without FCC pretreaters. Therefore,

rather than simply use the PADD average conversion of FCC feedstock to naphtha for all 

refineries in a given PADD, the refinery-by-refinery model differentiates between refineries that 

have an FCC feedstock pretreater and those that do not. We have also quantified the gasoline 

blendstock fraction yield for FCC units that have both feed pretreater and postreater units.

Historically, refiners have installed FCC feed pretreaters for economic reasons, as 

pretreaters increase FCC unit conversion to high value gasoline blendstock while decreasing the 

production of low value light cycle oils and residual material from FCC units. FCC feed 

pretreaters also have the benefit of reducing sulfur from the FCC feedstocks, resulting in the 

production of lower sulfur FCC naphtha and ultimately lower sulfur gasoline. In developing our 

refinery-by-refinery model, we quantified the impact FCC feed pretreating and postreating has 

on FCC naphtha yields and sulfur levels based on our evaluation of information we received 

from technology vendors. The results of this analysis are shown in Table 5-8 below.

Table 5-8 FCC Unit Gasoline Blendstock Fraction Yields

Average of All 

FCC Units

FCC Units with 

No Pretreater

FCC Units with a 

Pretreater Only

FCC Units with a 

Pretreater and 

Postreater

PADD 1 0.560 0.558 0.638 0.607

PADD 2 0.570 0.533 0.648 0.617

PADD 3 0.554 0.520 0.630 0.600

PADD 4/5 excluding 

California

0.565 0.548 0.642 0.612

5.1.3.1.2.1 Poly Gas and Alkylate

For the polymerization and alkylation units the capacity of the unit coupled with its 

estimated utilization rates listed in Table 5-5 is sufficient to establish the volume of gasoline 

blendstock produced by these units. For example, a particular refinery unit in PADD 1 might 

have a 10,000 barrel per day alkylation unit. If the alkylation units in PADD 1 are estimated to 
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

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be operating at 56 percent of its listed capacity in 2017, the alkylate production is projected to be 

5,600 barrels per day at that refinery. Each of the refineries within a given PADD was assumed 

to have the same utilization rate for any alkylation units.

5.1.3.1.2.2 Light Straight Run Naphtha

The remaining gasoline blendstocks, including light straight run naphtha (LSR), coker 

naphtha and hydrocrackate cannot be estimated simply using the unit capacity and unit 

utilization rate. In order to determine the volume of gasoline blendstock produced by each of 

these units, additional steps are required. LSR naphtha is principally comprised of five- and sixcarbon hydrocarbons which come directly from crude oil. Thus the volume of LSR for each 

refinery was based on the volume of crude oil processed by each refinery as determined in 

Section 5.1.3.1.1, as well as the percentage of that crude oil that is LSR. The fraction of LSR in 

each refinery’s crude oil was estimated on a PADD average basis using the LP refinery model

since it is not available on a refinery-by-refinery basis. This percentage is based on the types and 

quality of crude oil processed by all the refineries in each PADD from our LP model4. LSR as a 

percentage of crude oil is estimated to vary from 4 to 5 percent across the PADDs. These PADD 

level results are shown in Table 5-9 below.

Table 5-9 LSR as a Percentage of Crude Oil by PADD

PADD 1 PADD 2 PADD 3 PADDs 4/5a

LSR as a Percentage of Crude Oil 4.5% 5.0% 4.4% 4.4%

a Excluding California

After we calculated how much LSR is produced at each refinery we determined how 

much of the LSR is used as a gasoline blendstock. LSR has several possible destinations that 

vary from refinery to refinery. For each refinery, with the exception of those located in PADD 2, 

a portion of the LSR is designated to be sold into the petrochemicals market where it is 

processed into other hydrocarbon compounds. EIA publishes the volume of naphtha which is 

sold into the petrochemicals market in each PADD5. This information is summarized in Table 

5-10 below. Since this information is not publicly available on a refinery-by-refinery basis, we 

assumed that the volume of LSR naphtha sold into the petrochemicals market by each refinery is 

proportional to the refinery’s percentage of the total volume of crude oil processed in the PADD 

in which the refinery is located. After accounting for the volume of LSR naphtha sold to the 

petrochemicals market, the balance of LSR naphtha is used as a feedstock for the isomerization 

unit if the refinery has one. If a refinery does not have an isomerization unit, all of the LSR not 

sold to the petrochemical market is assumed to be used as a gasoline blendstock. Any volume of 

LSR at a given refinery that exceeds the capacity of the isomerization unit at the facility is also 

assumed to be used as a gasoline blendstock. However, if a refiner does not have enough naptha 

hydrotreating capacity to process all of the refiner’s LSR volume, we assumed that the refiner 

would use excess capacity in their FCC postreater, to reduce the sulfur content of the LSR 

blendstock. 
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Table 5-10 Refiner Sales of Naphtha in 2009 (1,000 BPSD)

PADD 1 PADD 2 PADD 3 PADD 4/5a

Naphtha Sold to the Petrochemical Industry 12.2 22.4 161.2 0

Sales of Special Naphtha 0.8 0 31.4 0

a Excluding California data

For further clarity on gasoline blendstock yields from the model, the gasoline blendstock 

volumes of LSR and naphtha from the naphtha splitter overhead tower are adjusted to subtract 

sales of these blendstocks that are sent to the petrochemicals market. E The values listed in Table 

5-11 for LSR and naphtha splitter overhead are the volumes sent to gasoline, as a fraction of 

crude throughput. 

Table 5-11 PADD Average Gasoline Blendstock Yields per Fraction of Crude Input

PADD 1 PADD 2 PADD 3 PADD’s 4/5a

LSR to Gasoline 0.0309 0.0286 0.0187 0.0381

Naphtha from Light Naphtha Splitter Overhead 0.0021 0.0081 0.0072 0.0093

a Excluding California

In refineries with an isomerization unit, much of the LSR is processed into isomerate, the 

product produced by the isomerization unit. The volume of isomerate produced is dependent on 

the volume of feedstock processed by the isomerization unit up to its capacity. As described 

above, all of the LSR that is not assumed to be sold into the petrochemical markets is assumed to 

be sent to the isomerization unit, up to the maximum capacity of the isomerization unit. The 

isomerization unit produces a blendstock with a slightly higher energy density and smaller 

volume compared to the feedstock volume. To account for this effect, the volume of isomerate 

produced is estimated to be 1.6 percent less than the volume of LSR feedstock to the 

isomerization unit. Hydrocrackate and Coker Naphtha

The hydrocracker and coker units also produce some light naphtha material which is 

blended into gasoline. Heavy naphtha is also produced in these units, which is feed to the 

reformer, as discussed in the next section. The light naphtha material produced by the 

hydrocracker and coker are termed light hydrocrackate and light coker naphtha, respectively. 

Based on LP refinery modeling work done for the MSAT2 rule we estimated that the portion of 

the feedstock processed by each of these units converted to light coker naphtha and light 

hydrocrackate was 5 percent for coker units across all the PADDs, and ranges from 23 to 32 

percent for hydrocracker units depending on the PADD. The light coker naphtha is poor in 

quality and require hydrotreating to removes sulfur, olefins and other impurities, before sending 

them to an isomerization unit, if a refiner has one. Table 5-12 below summarizes the percentage 

of the feedstock to these units that is converted to light naphtha and blended into gasoline.

 

E Naphtha Splitter towers separate the naphtha feed stream into a light and heavy streams, whereby the heavy stream 

is typically reformer feedstock, while the light stream is blend stock lighter than reformer feed. 
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Table 5-12 Isomerization Unit Feed Rates by PADD

PADD 1 PADD 2 PADD3 PADD 4/5a

Light Coker Naphtha

(% of Coker Feed)

5.0% 5.0% 5.0% 5.0%

Light Hydrocrackate 

(% of Hydrocracker Feed)

28.7% 32.0% 23.3% 27.2%

a Excluding California

5.1.3.1.2.3 Reformate

The volume of reformate produced by the reformer was estimated based on the volume of 

feed to the reformer as limited by each unit’s capacity. The feed to the reformer comes from 

various sources depending on the refinery configuration. For virtually all refineries, the heavy

part of the straight run naphtha from the atmospheric crude tower is sent to the reformer, while 

the light naphtha is generally processed in the isomerization unit or blended directly into gasoline 

as discussed above. Those refineries with a hydrocracker or a coker will send the heavy naphtha 

from these units to the reformer as well. This reformate feed naphtha contains the six, seven, 

eight and usually the nine carbon compounds from these various sources. In some cases, the six 

carbon compounds are separated from the rest of the reformate feedstock to reduce the benzene 

in the final reformate. The volume of the feed to the reformer is estimated based on a fractions 

of the material processed in the atmospheric crude tower, hydrocracker and coker on a PADD by 

PADD basis using information from the LP refinery model. 

The fraction of crude oil that is fed to the reformer from the atmospheric crude tower 

ranges from about 13 to 16 percent of the crude oil input depending on the PADD. About 18 

percent of the material processed in the coker unit is estimated to end up as feedstock to the 

reformer. The percentage of the feedstock processed in the hydrocracker that is fed to the 

reformer ranges from 30 to 50 percent depending on the PADD in which the refinery is located. 

The variance in the fraction of hydrocracker material sent to the reformer is due to the significant 

flexibility that the hydrocracker has for producing either gasoline or diesel fuel. In certain 

PADDs, such as PADD 4 and 5, there is a higher relative demand for diesel fuel compared to 

gasoline so there is a lower conversion to naphtha than in other PADDs. The product from the 

reformer experiences a volume decrease of about 18 percent relative to the volume of feed, due 

to the conversion of straight chain and cyclical hydrocarbons to energy dense aromatics and 

other light products. This volume reduction and conversion to lighter products increases with the 

severity and thus the conversion of the reformer unit. All the refineries in each PADD are 

assumed to be operating their reformers at the same severity as estimated by the LP refinery

model. Each of the values discussed in this paragraph are shown on a PADD by PADD basis in 

Table 5-13 below.

Table 5-13 Reformer Feed Rates and Volume Loss

PADD 1 PADD 2 PADD3 PADD 4/5a

Medium/Heavy Straight Run Naphtha 

(% of Crude Input)

13.8% 16.2% 14.0% 13.6%

Medium/Heavy Coker Naphtha 

(% of Coker Feed)

18.4% 18.4% 18.4% 18.4%
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Medium/Heavy Hydrocrackate 

(% of Hydrocracker Feed)

35.4% 43.4% 50.2% 33.3%

Volume Loss in Reformer 18% 17% 18% 19%

a Excluding California

5.1.3.1.2.4 Purchased Blendstocks

Some gasoline blendstocks are purchased and blended into gasoline. The gasoline 

blendstocks typically purchased include natural gasoline, alkylate, isooctene and ethanol. We 

did not have information on the volume of these gasoline blendstocks purchased and blended 

into gasoline by each refinery, so we again relied on the information from EIA, which reports the 

consumption of these blendstocks on a PADD basis. The EIA information on the amount of 

pentane plus, naphtha’s and NGLs purchased in each PADD are listed in Table 5-14 below. Our 

RFG database has each refiners amount of ethanol blended into RFG, but does not contain the 

amount of ethanol that is splash blended into CG at terminals. We accounted for ethanol blended 

into CG, as well as the purchase of other gasoline blendstocks, by assuming that each refinery

purchased a volume of any given gasoline blendstock purchased within their respective PADD 

proportional to that refinery’s crude oil consumption within the PADD. In the 2009 RFG 

database, the ethanol volumes only averaged 2.7 percent of refiner’s gasoline production, which 

results in an over estimation of our refinery and program costs in this NPRM. In our NPRM 

analysis, we did not include any desulfurization costs for Pentane plus and naphtha and lighter 

blendstocks, since we did not know the extent that they were being treated today. However, the 

addition of these blendstocks, results in very negligible increases in demand for additional 

naphtha hydrotreating. We will evaluate whether we need to include any costs for treating these 

streams in the final rule analysis. If there are costs, because the streams are so small, the costs 

would be negligible. 

Table 5-14 Refiner Purchases in 2009 (1,000 BPSD)c

PADD 1 PADD 2 PADD 3 PADD 4/5a

Natural Gas Liquids 17.5 115.9 272.4 15.5

Naphtha’s and lighter 24.6 39.3 53.7 3.8

Pentanes Plus 0 41.8 94.9 25.7

Ethanol b

57.3 66.9 77.3 0.40

Notes:

a Excluding California

b Ethanol from EPA RFG database, excluding volumes that are splash blended into conventional gasoline

c Natural Gas Liquids and Pentanes Plus are different names for the same hydrocarbon stream and we 

inadvertently found two different volumes for the same hydrocarbon stream and added them both as inputs into 

our refinery cost model. We will correct this in the final rulemaking analysis.Butane Volumes

To estimate the butane volumes in our refinery-by-refinery model we used an RVP 

balance equation. This equation states that the product of the overall RVP and volume of the 

gasoline pool is equal to the sum of the product of the RVP and volumes of the non-butane 

components plus the product of the RVP and the volume of the butane blendstocks. This 

equation can be rearranged to solve for the volume of butane blendstocks as shown in Equation 

5-1 below. 
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Equation 5-1 RVP Butane Balance Equation

Butane = (A*D-B*D)/(C-A)

Where:

Butane = Volume of Butane added in each refinery in BPSD

A = Blended gasoline RVP average

B = Non-butane blendstock RVP average

C = Butane RVP

D = Volume of gasoline produced

The gasoline production volumes and RVP of the blended gasoline are reported to EPA 

by refiners for each refinery and were used for the A and D terms in Equation 5-1. To calculate 

the RVP of the butane used as gasoline blendstock we first had to consider the relative 

proportion of isobutane versus n-butane being used as a gasoline blendstock as their RVP values 

differ. This ratio was estimated on a PADD by PADD basis from the LP modeling work. We 

then used a volume weighted average to calculate the RVP of the mixed butane stream blended 

into gasoline in each PADD. The information for these calculations is shown in Table 5-15

below. The non-butane blendstock RVP was estimated by multiplying each individual gasoline 

blendstock RVP times the gasoline blendstocks volume fraction of each refineries gasoline pool 

(CG and RFG) using 2009 ethanol volumes and taking the sum of all of these values. The RVP 

value for each of these streams is shown in Table 5-16 below. With this information we were 

then able to estimate the volume of butane added to the gasoline blendstock at each refinery. 

The annual volumes of butane added by refineries on a PADD level are listed in Table 5-17. The 

volume of butane blended into gasoline at each individual refinery varies based on the annual 

average gasoline RVP that the refinery produces ( the RVP of CG and RFG gasoline are volume 

weighted together), as well the variance in gasoline blendstock streams that a particular refinery

uses to produce CG and RFG gasoline.

Table 5-15 PADD Average Composition of Mixed Butanes Added to Gasoline

PADD 1 PADD 2 PADD 3 PADD’s 4/5a

Isobutane % 96% 32% 53% 66%

N-butane % 4% 68% 47% 34%

Mixed Butane RVP (C), psi 71.376 58.192 62.518 65.196

a Excluding California
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Table 5-16 PADD Average RVP’s of Gasoline Blendstocks

PADD 1 PADD 2 PADD 3 PADD’s 4/5a

LSR 12.0 12.0 12.0 12.0

Naphtha from Light Naphtha 

Splitter Overhead

3.0 3.0 3.0 3.0

Reformate 4.5 6.6 5.0 6.2

FCC Naphtha 4.6 4.6 4.6 4.6

Coker Naphtha 13.0 13.0 13.0 13.0

Isomerate C5 13.0 13.0 13.0 13.0

Isomerate C6 7.2 7.2 7.2 7.2

Natural Gasoline (NGL) 12.6 12.6 12.6 12.6

Polymerization Gasoline 2.8 2.8 2.8 2.8

Light Hydrocrackate 9.2 9.2 9.2 9.2

Alkylate, C3 3.6 3.6 3.6 3.6

Alkylate, C4 3.2 3.2 3.2 3.2

Dimersol 5.8 5.8 5.8 5.8

Ethanol 10.7 10.7 10.7 10.7

a Excluding California data

Table 5-17 PADD Average Gasoline Data

PADD 1 PADD 2 PADD 3 PADD’s 4/5a

Non-butane blendstock RVP (B), psi 6.4 6.3 5.7 6.3

Gasoline Pool Volume (D), BPSD 738.2 1744 3487.8 500.3

Volume Butane Added, BPSD 21.4 95.1 143.1 25.7

Blended Gasoline RVP average (A), psi 8.5 9 8 9.1

a Excluding California data

5.1.3.2 Calibrating the Blendstock Volumes in the Refinery-By-Refinery Model

After calculating gasoline volume estimates for each refinery in the refinery-by-refinery

cost model, we calibrated these values against their reported gasoline blendstock volumes. 

Refiners report their production volumes for both conventional and reformulated gasoline to 

EPA to comply with the gasoline reporting requirements. We used these reported volumes from 

2009, along with LP modeling results from our MSAT2 Rule, to calibrate the refinery-byrefinery model. Before making any adjustments, the refinery-by-refinery modeling estimates for 

gasoline produced on a national basis correlated very well with the reported refinery production 

volume in 2009, with volumes differing by less than 0.5 percent from actual production. In order 

to eliminate this discrepancy we modified each refiner’s yields in our refinery-by-refinery

analysis based on the 2009 data as reported by the refineries. The volume of each of the gasoline 

blendstocks, excluding the light straight run (LSR) and FCC gasoline streams, were increased or 

decreased proportionally in order to align the aggregated national finished gasoline production 

volumes in our refinery-by-refinery model with the aggregated national finished gasoline 

production volume reported by US refineries in 2009.
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In making adjustments to the refinery-by-refinery analysis to better align its volumes with 

the reported gasoline volumes and sulfur levels, we did not make any changes to the production 

volumes of LSR or FCC gasoline. The LSR and FCC gasoline volumes are left unchanged as 

these volumes are based on actual refinery-specific FCC charge throughput rates. This accounts

for the effects of FCC feed pretreating on FCC gasoline yields and is therefore likely to 

accurately reflect the production volumes for each individual refinery. Additionally, these 

volumes are of central importance to our analysis as they are used for estimating the equipment 

modifications necessary for complying with new Tier 3 sulfur standards and the costs associated 

with additional sulfur removal. 

With these calibrated volumes for each of the gasoline blendstocks, the refinery-byrefinery model can now be used to estimate the sulfur level that refiners must achieve in the FCC 

naphtha to meet the current sulfur limit under the Tier 2 standards. These volumes also allow us 

to model what modifications to existing equipment and refinery operations will be required to 

comply with the new Tier 3 sulfur standards

5.1.3.3 Refinery Blendstock Sulfur Levels

After determining the volume of each gasoline blendstock stream, we next estimated the 

sulfur level of each of the gasoline blendstocks for our modeling analysis using information we 

collected from literature reviews and discussions with refinery consultants and technology

providers. We also considered the blendstock sulfur levels estimated for the MSAT2 rule and 

the estimates derived from our refinery-by-refinery model to estimate the sulfur levels of the 

blendstock streams. Establishing these sulfur levels is important as this sets a baseline for the 

refinery-by-refinery model that represents our estimate for the current operations of each 

refinery. This allows us to project what changes refiners would have to make in their refineries 

to comply with the Tier 3 standards, and project the new investments and operating costs 

associated with these changes. The following section contains further details on how the sulfur 

content of each of the blendstocks was estimated. The results of this analysis can be found in 

Table 5-18 at the end of this section.

The first stream we considered was the butanes that are used as a gasoline blendstock. 

The butanes used as gasoline blendstock within a refinery come from a variety of sources. Much 

of the butane used as a gasoline blendstock is distilled from the crude oil or other blendstock 

streams within the refinery. Refiners remove the butanes from crude oil and sometimes gasoline 

blendstocks which contain some butane (i.e., FCC naphtha, hydrocrackate) and then blend them 

back into the gasoline up to the RVP or vapor/liquid limit applicable to the gasoline market that 

the gasoline is being sold into. During the summer months refiners usually have excess butane 

which cannot be blended into the gasoline pool because of the tighter RVP standards. Many 

refiners store the excess butanes and then blend them back into gasoline in the winter months 

when the volatility limits for gasoline are less stringent. Other sources of butanes used as 

gasoline blendstocks are natural gas processers and crude oil drilling operations. The butanes 

from these sources are produced in downstream units which separate the various hydrocarbon 

components. Most of these downstream units “sweeten” the butanes using a Merox unit prior to 

shipping them in pipelines or selling them directly to refiners. The sweetening process reacts the 

hydrocarbon mercaptan compounds to disulfide compounds reducing their odor and corrosivity. 

The sweetening process, however, does not lower the sulfur level. If the source natural gas well 
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-21

is very high in sulfur, the operator may need to use an extractive Merox treatment technology

which actually removes the sulfur from the butane stream. This treatment generally lowers the 

sulfur level of the butanes to under 5ppm. Butanes that are blended into gasoline have a sulfur 

limit of 30 ppm and those that are shipped through pipelines, regardless of their end use, have a 

limit of 140 ppm. Furthermore, many refiners have Merox units on site that are capable of 

removing sulfur from butanes that are either purchased or generated internally from refinery

units. We were, however, unable to evaluate existing butane Merox treating capacity at NGL 

processers, crude drilling operations, or in refineries as there was no information available in the 

OGJ, from EIA, or other publically available sources. Because we do not know the prevalence 

of these units, we conservatively assumed in our baseline case that refiners are adding treated 

butanes with a sulfur content of 10 ppm to their gasoline pool.F 

For hydrocrackate, dimersol, and poly gas blendstock streams, we used the same sulfur 

levels that we estimatd for our MSAT2 rulemaking. The sulfur levels for these streams are 

inherently low due to the dynamics of process reactions in the hydrocracker, dimersol and 

polymerization units. Furthermore, it is unlikely that refiners have altered these processes in 

their refineries since our analysis for the MSAT2 rule was completed.

Alkylate blendstocks usually have a small amount of sulfur, usually less than 15 ppm. 

The primary source of sulfur in alkylate is the sulfuric acid that is used as a catalyst in the 

alkylation process. Finished product coalescers and knockout drums are used by refiners to 

remove impurities, including sulfuric compounds, from the alkylate product as it leaves the 

alkylation unit. This separation is imperfect, and a small quantity of the sulfuric compounds 

which remain in the alkylate account for the majority of its sulfur content. Prior to the enactment 

of the Tier 2 standards, the alkylate produced by most refineries contained 10 to 15 ppm sulfur 

which assumes that there was some carryover of sulfuric compounds into the alkylate. Based on 

our discussions with gasoline desulfurization technology vendors, however, refiners have 

installed new acid coalescers and knock out drums in recent years. These new units improved 

the removal of residual sulfuric compounds and can produce an alkylate blendstock with a 5-ppm 

sulfur level. This adjustment by refiners seems to be a low cost method for reducing the sulfur 

content of alkylate. For our refinery-by-refinery baseline analysis, we assumed that refiners have 

already installed improved acid knockout drums and are currently producing a 5-ppm alkylate. 

We also assumed that Hydrofluoric Acid (HF) alkylation processes had the same alkylate yield

per feedstock throughput as a sulfuric acid alkylation unit in our refinery by refinery model. We 

assumed that the sulfur level of alkylate from an HF units also averages 5 ppm sulfur, even 

though HF processing units use hydrofluoric acid as the processing catalyst, instead of using

sulfuric acid. 

The coker unit produces a gasoline blendstock with a significant amount of sulfur. These 

units convert the heavy portion of crude oil, called residuals, into gasoline and diesel blendstocks 

through the use of heat and pressure. The gasoline blendstock produced by the coker can contain 

 

F After we completed our cost analysis, we met with UOP staff, including those who market their Merox technology 

for removing mercaptans from gasoline streams. The UOP staff said that pretty much all butanes are already being 

treated by Merox (or similar) extraction units. Thus, there would be no additional cost for treating butanes for 

complying with Tier 3. We will update our cost analysis to reflect this for the final rule cost analysis. 
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more than 3,000 ppm sulfur. This stream is normally split into two different streams. The 

stream which contains the six to nine carbon hydrocarbons is processed in the naphtha 

hydrotreater, which reduces the sulfur level of this blendstock to below 1 ppm. This stream is 

then routed to the reformer for octane improvement. The five and six carbon hydrocarbon 

portion of coker naphtha is called light coker naphtha and usually contains on the order of 

several hundred ppm sulfur. Because of the instability of this stream due to its high olefin 

content, it is generally processed by the naphtha hydrotreater and sent to the isomerization unit if 

the refinery has one. After being processed in the hydrotreater, the sulfur content of this stream 

is reduced to approximately 1 ppm. These treating pathways were assumed for each refinery in 

the refinery-by-refinery baseline analysis.

Straight run naphtha is a gasoline blendstock which contains a moderate amount of 

sulfur. Straight run naphtha is the product stream from the atmospheric crude oil tower with a 

boiling point that falls within the boiling range of gasoline. The heaviest portion of straight run 

naphtha is higher in sulfur relative to the lighter portion of the straight run naphtha. The heavy 

portion of straight run naphtha is normally processed by the reformer in order to improve its 

octane before being blended into gasoline. After this processing, the reformate has a sulfur level 

of less than 1 ppm. The light straight run naphtha (LSR) contains the five and part of the six 

carbon hydrocarbons and has on the order of 100 ppm sulfur before any hydrotreating. LSR that 

is routed as feedstock to isomerization units has its sulfur lowered to 1 ppm by processing in the 

naphtha hydrotreater. This hydrotreating is necessary to allow this material to be processed in 

the isomerization unit, as the catalysts in these units require low sulfur feedstocks to function 

properly. Some refiners, however, do not have isomerization units or they produce LSR volumes 

that are greater than the capacity of their isomerization units. Even cases where there is 

insufficient capacity in the isomerization units it is still desirable for refiners to hydrotreat as 

much of the LSR as possible since it is more cost-effective to reduce the sulfur content of the 

LSR than the FCC naphtha. Refiners can either hydrotreat this volume of LSR in the naphtha 

hydrotreaters or in FCC naphtha postreaters.

Natural Gas Liquids (NGL) have a composition that is similar to LSR, as it is comprised 

primarily of pentanes and hexanes. NGLs are produced from natural gas processers and crude 

oil drilling operations and the sulfur content of the NGLs can vary depending on its source, 

although we estimate that this stream averages about 100. While some of the NGLs are treated 

to remove sulfur by the NGL producers before being purchased by the refineries we did not have 

sufficient information to be able to determine the extent to which NGLs are treated before 

arriving at the refinery. For the baseline case in our refinery-by-refinery model we assumed that 

NGL liquids are purchased with a sulfur content of 100 ppm and hydrotreated based on capacity

availability at refineries in a similar manner as LSR. Based on the gasoline blendstock volumes 

and hydrotreating capacities as discussed in the previous sections, we estimated in the baseline 

case for our refinery-by-refinery analysis that refiners are hydrotreating 66 percent of the 

volumes of LSR and NGLs produced and purchased for gasoline blendstock usage on a national 

average basis. Our hydrotreating capacity evaluation for each refinery is discussed in more 

detail in Section 5.1.3.4. As a result of the proposed sulfur standards under the Tier 3 program,

we anticipate that refiners will revamp existing hydrotreaters and add new hydrotreating capacity 

to allow them to hydrotreat all of their LSR and NGL material. 
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5-23

We also assumed that all ethanol blended into gasoline has a sulfur content of 5 ppm. 

Ethanol produced at ethanol plants naturally has a negligible amount of sulfur. Before being 

sold, however, a denaturant is added to the ethanol. This denaturant most commonly used is 

natural gasoline, a C5 to C7 naphtha produced during natural gas processing. Natural gasoline 

has a sulfur content that ranges anywhere from a few parts per million to a couple hundred parts 

per million sulfur. We assumed that the natural gasoline used as an ethanol denaturant is not 

hydrotreated and has an average sulfur level of 250 ppm. Ethanol contains 2 percent denaturant, 

which results in denatured ethanol having a sulfur level of 5 ppm.

After determining the sulfur level for each of the gasoline blendstock streams as 

discussed above we can use this information, along with the gasoline production volumes and 

sulfur levels for the United States in 2009, to determine the sulfur level of the FCC naphtha 

stream on a national average basis. To do this we used the following equation, referred to as 

Equation 5-2 hereafter:

FCC Naphtha Sulfur ppm = [(A*B) – (C*D+E*F+G*H+I*J+K*L+M*N+O*P+Q*R+S*T)] / Z

Where:

A = Refinery Total Gasoline Yield, BPSD

B = Refinery Total Gasoline Sulfur level, ppm

C = Butane to Gasoline, BPSD

D = Butane Sulfur, ppm

E = Alkylate BPSD

F = Alkylate Sulfur, ppm

G= Reformate BPSD

H= Reformate Sulfur, ppm

I = Coker Naphtha, BPSD

J = Coker Naphtha Sulfur, ppm

K= Hydro-crackate BPSD

L= Hydro-crackate Sulfur, ppm

M= Light Straight Run (LSR) and Natural Gas Liquids (NGL), BPSD

N =LSR and NGL Sulfur, ppm

O= Dimersol, BPSD

P= DimersolSulfur, ppm

Q= Polymerization BPSD

R= Polymerization Sulfur, ppm

S= Ethanol, BPSD

T = Ethanol Sulfur, ppm

Z= FCC Gasoline Yield, BPSD

Equation 5-2 Calculating FCC Naphtha Sulfur Content for Refinery-By-Refinery Model

We used this equation to assess two cases; a baseline case where the 30-ppm Tier 2 sulfur 

standards were fully implemented and a control case that reflects the proposed 10-ppm Tier 3 

sulfur standards. The only terms in Equation 5-2 that change between the two cases are the 

national average sulfur level and the sulfur levels of the LSR, NGL, and FCC naphtha streams. 

The national average sulfur levels for the two cases were set at the sulfur limits under the Tier 2 
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and Tier 3 programs -- 30 ppm and 10 ppm, respectively. For the baseline case we assumed that 

the sulfur level of the NGL and LSR streams was 34 ppm. This reflects our assessment of how 

these streams are currently being handled as discussed earlier in this section. We estimate that 

66 percent of the volume of NGL and LSR are hydrotreated before being blended into gasoline 

and have a very low sulfur content of approximately 1 ppm. The remaining 34 percent are 

untreated and have a sulfur content of approximately 100 ppm. For the Tier 3 control case we 

assumed that all of the NGLs and LSR were hydrotreated and therefore had an average sulfur 

content of 1 ppm. This information allowed us to solve Equation 5-2 for the FCC naphtha 

content. The resulting FCC naphtha sulfur numbers, along with our estimation of the gasoline 

blendstock sulfur levels and percent of total gasoline volume made up by each blendstock are 

shown in Table 5-18 below.

Table 5-18 Sulfur Levels for Gasoline Blendstocks in the Refinery-By-Refinery Model

Gasoline Blendstocks Baseline Tier 2 

Case

Proposed Tier 3 

Case Year 2017

Proposed Tier 3 

Case Year 2030

Percent 

of Total 

Volume

Sulfur 

Levels 

30 

ppm

Percent 

of Total 

Volume

Sulfur 

Levels 

10 

ppm

Percent 

of Total 

Volume

Sulfur 

Levels 

10 

ppm

FCC Naphtha 37.2 80a36.0 21

a

35.0 21

a

Reformate 22.5 0.5 21.8 0.5 21.2 0.5

Alkylate 12.7 5 12.5 5 12.1 5

Isomerate 3.2 0.5 3.1 0.5 3.1 0.5

Butane 4.0 10 4.0 10 3.8 10

Light Straight Run Naphtha (LSR) and 

Natural gas Liquids (NGL)

5.2 34 4.9 1 4.8 1

Hydrocrackate 3.0 8 2.9 8 2.8 8

Ethanol 9.9 5 12.5 5 15 5

Coker Naphtha 2.2 0.5 2.1 0.5 2.0 0.5

Other Gasoline Blendstocks 0.2 10 0.2 10 0.2 10

Total/Sulfur Average 100 30 100 10 100 10

a These values are calculated using Equation 5-2; all other sulfur levels are assumed

The numbers in the table above represent national averages. While this is useful 

information, it is insufficient for us to be able to model the implications of the proposed Tier 3 

standards for an individual refinery. Each refinery has a unique combination of processing units 

that will determine the cost and operational changes necessary for that refiner to comply with our 

proposed sulfur limit. While each of these processing units may impact the cost for refiners to 

lower the sulfur content of the gasoline they produce we believe these costs will be dominated by

the units responsible for the desulfurization of the FCC naphtha, and to a lesser extent the NGLs 

and LSR. This is because these are the only streams we anticipate would see significant sulfur 

reduction under the proposed Tier 3 sulfur standards. The units that are used to desulfurize these 

streams include the FCC unit pre- and postreaters and the naphtha hydrotreaters. It is important, 
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therefore, to have a good understanding of which of these units are in place in each refinery, as 

well as the type and capacity of these units, in order to allow us to most accurately estimate the 

cost of the Tier 3 sulfur standards to the refining industry. We used the above FCC naphtha 

sulfur balance information as the basis of our vendor request for refiner modifications to FCC 

postreaters under Tier 3. However, for the vendor requests, we used a preliminary model, where 

the FCC naphtha levels under Tier 2 averaged 75 ppm, while FCC naphtha levels under Tier 3 

averaged 25 ppm for 10-ppm sulfur gasoline, representing a 50 ppm sulfur reduction, close to the

same delta presented in the table above. The following section discusses our assessment of the 

desulfurization equipment currently being used in refineries.

5.1.3.4 Assessment of Refineries’ Existing Desulfurization Equipment

Since the desulfurization cost of the Tier 3 program is largely impacted by the cost of 

lowering sulfur in FCC gasoline, it is important to understand what refiners are already doing to 

lower the sulfur content of the FCC gasoline blendstock to meet the Tier 2 sulfur standards. This 

was important to our analysis of the cost for each individual refiner to reduce the sulfur content 

of their gasoline to meet the proposed Tier 3 sulfur standard. Refiners that already have an FCC 

pretreater or postreater can revamp these units for a lower cost than installing grass roots units. 

It was also important to determine which refineries have an FCC feed pretreater, since these units 

increase the refineries FCC conversion and production of FCC naphtha and also lower the sulfur 

level of the FCC naphtha. To compile this information we analyzed capacity information for 

FCC naphtha pretreaters and postreaters for each refinery listed in the OGJ and the EIA database. 

If one of the databases showed that a refinery had FCC pretreating and/or post-treating capacity,

while the other did not, we assumed that the refinery did have the units listed with a capacity as 

reported.

For refineries that have FCC naphtha postreaters we next determined which vendor’s 

FCC naphtha desulfurization technology is installed in each refinery. To do this we conducted a 

public database search using OGJ, company web postings and, other refinery publications. To 

supplement this data we also had extensive discussions with many refiners to obtain confidential 

data from many of them on type and capacity of the desulfurization technology currently 

installed in their refineries as well as how their operations might be adjusted to meet the new 

Tier 3 sulfur standards. The various FCC naphtha desulfurization technologies that we identified 

as currently being used by refiners are CD Tech’s Cd Hydro and CDHDS, Axens Prime G and 

Prime G+, UOP’s ISAL and Selectfining, Exxon’s Scanfining I or II and Sinopec’s S-Zorb. For 

refiners that we could not find or obtain information on the type of desulfurization they were 

using, Axens was chosen as the default as they have the largest market share of desulfurization 

units in the U.S. To confirm the accuracy of our work we reviewed our assessments with one of 

the main technology vendors. Our desulfurization technology selection assumptions were 

adjusted based on feedback from the vendor. The aggregated results of this assessment are 

summarized in Table 5-19 below.

Table 5-19 Postreater Technologies Used By Refineries

CD Tech IFP/Axens Scanfining UOP ISAL S-Zorb

Refiners with Existing 

Postreater

15 40 16 1 4
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The next step of our analysis was to determine which refineries use FCC feed 

hydrotreating technology (pretreaters) in addition to post-treating units. FCC feed hydroteating 

was primarily installed at refineries not as a sulfur control technology, but because of the 

economic benefits that can be obtained from hydro-treating FCC feed. Hydrotreating the FCC 

feed increases the crackability of this stream by saturating the components with hydrogen 

resulting in a higher paraffin content in the feed stream. Hydrotreating also removes FCC feed 

impurities such as nitrogen, metals, con-carbon and sulfur, which improve FCC unit catalyst 

effects. An additional benefit of FCC feed pretreating is that it reduces the sulfur content of the 

FCC feedstock by 70 to 90 percent, resulting in the production of FCC naphtha with lower sulfur 

levels than what would be produced using FCC feed that is not hydrotreated. 

Our analysis indicates that approximately 53 refiners are currently using FCC feed 

pretreaters. Of the 53 refineries with pretreaters, 35 of also have FCC postreaters installed to 

comply with the Tier 2 gasoline sulfur standard. The technologies used by these 35 refineries are 

shown in Table 5-20 below. FCC naphtha produced using only an FCC pretreater operating at 

standard severity generally produces a gasoline with a sulfur content that exceeds the Tier 2 

standards. According to information from vendors, the average FCC naphtha sulfur level of 

refineries with an FCC feed pretreater operating at standard conditions without a postreater

ranges from 200 to 500 ppm. Further reductions in the sulfur level of the FCC naphtha are 

possible using only an FCC pretreater by operating the pretreater at a higher severity or higher 

pressure (if the unit is designed to do so). These high pressure FCC pretreating units were 

designed to be able to run at a high severity to further increase the crackability of the FCC feed 

and therefore increase the conversion rate of the FCC unit. These more severe conditions also

further reduce the sulfur level of the FCC naphtha. The naphtha produced from these units 

operating with high severity or high pressure has an average sulfur content ranging from 75 to 

100 ppm, allowing these refineries to produce gasoline that meets the Tier 2 sulfur standards. 

Operating FCC pretreaters at the high severities necessary to meet the Tier 2 standards, however, 

also results in increased operating cost, as the pretreater requires more frequent catalyst 

changeouts, consumes more hydrogen, and operates higher temperatures than pretreaters 

operating under standard conditions.

Table 5-20 Technologies Used By Refiners with FCC Pre and Postreaters

CD Tech IFP/Axens Scanfining UOP ISAL S-Zorb

Refiners with FCC Pretreater 

and Naphtha Postreater

9 20 5 0 1

Our analysis also showed that there are several refineries that have an FCC unit but have 

installed neither an FCC naphtha postreater nor an FCC feed pretreater. These are small 

refineries, or refineries that produce a refinery gate gasoline with a sulfur level below the Tier 2 

cap of 80 ppm sulfur, but above the 30-ppm average. These refiners are relying on buying or 

sharing sulfur credits from other refineries that are over-complying with Tier 2 and make 

gasoline with a sulfur level less than 30 ppm. 

Finally, some refineries do not have an FCC unit and therefore have not installed FCC 

postreaters to comply with the Tier 2 sulfur standards. These refiners primarily use reformate, 

alkylate, LSR, butanes, and pentanes to make gasoline. Since these blendstocks all have low 
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sulfur content this allows refiners to produce gasoline with a low enough sulfur content to meet 

the Tier 2 sulfur standards. 

A summary of the number refineries which fall into differing categories of how they are 

complying with Tier 2 is shown in Table 5-21 below.

Table 5-21 Refinery FCC Naphtha Desulfurization Unit Characterization

5.1.3.5 Crude Oil and FCC Feed Sulfur Levels

After we had determined the desulfurization technology in place at each refinery, we 

sought to calculate the sulfur content of the FCC feedstock. Knowing this is important as it 

allows us to determine how far the sulfur level of the FCC naphtha, and ultimately the gasoline, 

produced at any given refinery can be reduced using the units currently in place at each refinery. 

It also helps us understand the extent to which the existing hydrotreaters are being taxed to 

comply with the Tier 2 gasoline sulfur standard. Some refineries may have excess capacity in 

their FCC naphtha pretreater or postreaters that would allow them to produce gasoline that would 

meet the proposed Tier 3 standards without having to revamp existing units or add grass roots 

units. These refineries will have much lower cost impacts than refineries that have to make more 

significant capital investments.

The sulfur level of the FCC feedstock is primarily dependent on the sulfur level of the 

crude oil being processed by the refinery and whether or not the refinery has an FCC feed 

pretreater. The first step, therefore, in determining the sulfur level of the FCC feedstock was to 

input the crude sulfur level for each refinery into our refinery-by-refinery model. For this, we 

obtained confidential business information (CBI) from EIA on the annual average crude sulfur 

levels that each refinery processed in 2009. This data, which is reported to EIA for each refinery, 

was used as the baseline crude sulfur level in our refinery-by-refinery analysis. Using this data, 

we then determined what each refiners FCC feed charge rate sulfur level would be, using a 

regression co-relation we built from data on crude sulfur levels and FCC feedstock material, as 

discussed below. We assumed that refineries today are primarily processing heavy gas oils 

(HGO) and vacuum gas oils (VGO) produced from each refinery’s crude and coker units. To 

determine the volume of feedstock processed by the FCC units we assumed that after distillation 

HGO makes up 20.5 percent (by volume) of the processed crude and VGO makes up an 

additional 15.6 percent of the crude.

6

Together, these two streams comprise the FCC feed. 

The boiling point range that we assumed for VGO also contained some residual material, 

representing FCC feed with residual content. This was done to reflect the imperfect distillation 

cuts in crude towers and that some refiners use small amounts of residual material as FCC 

FCC Treatment Units Installed Number of Refineries

No FCC Unit 14

FCC Unit, No Pretreater or Postreater 7

FCC Unit With Postreater Only 38

FCC Unit With Pretreater Only 17

FCC Unit With Pretreater and Postreater 35
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feedstock. The balance of the residual material, however, was excluded from the feed to FCCs 

since this material makes a poor feedstock due to its high aromatics, metals and concarbon 

content. Each of these materials negatively affects the FCC gasoline conversion yields. Most 

refiners today do not directly use residuals as feedstock to their FCC units, but instead send them 

to be processed in coker units or use the residual material for fuel oil and asphalt production. 

The boiling point ranges that we used for HGO and VGO are listed Table 5-22 below.

Table 5-22 Boiling Ranges of FCC Feedstocks

TBP Initial TBP Final

Heavy Gas Oil (HGO) 600°F 800°F

Vacuum Gas Oil (VGO)a800°F 1,000°F

a Contains some residual material

For our FCC feed sulfur regression, we used data obtained various crude oil assays that 

we obtained from Jacobs Engineering for work that Jacobs conducted for us. We used data 

from five specific crude types, including West Texas intermediate ,Bonny LightSaudi Heavy, 

Alaskan North Slope, and Mayan, and three blended crude assays. The equation for this 

regression, along with the estimated FCC feed sulfur contents for various crude oils are shown in 

Equation 5-3 and Table 5-23 below.

Equation 5-3 FCC Feed Sulfur Content Based on Crude Sulfur Content

FCC Feed Sulfur Weight Percent = (Crude Sulfur Weight Percent)0.8 * 1.1858 + 0.0409

Table 5-23 FCC Feed Sulfur For Various Crude Sulfur Levels

Crude Sulfur Level (Weight %) FCC Feed Sulfur Level (Weight %)

0.11% 0.24%

0.28% 0.47%

0.85% 1.08%

1.3% 1.50%

1.4% 1.59%

1.6% 1.77%

2.8% 2.74%

3.04% 2.93%

5.1.3.6 Impact of FCC Feed Pretreaters on FCC Feed Sulfur Levels

With FCC feed sulfur estimated, the next step in our analysis was to consider the impact 

of pretreaters on the FCC feed sulfur levels for refineries that have these units. There are several 

factors that must be considered to determine the impact of the pretreaters on the FCC sulfur 

level, including the pressure at which the unit operates, the severity at which it is run, and 

whether or not the FCC naphtha will be postreated. 

To inform our understanding of how FCC pretreaters operate, we obtained guidance from 

technology and catalyst providers. From these discussions we learned that the capability for 

FCC pretreaters to remove sulfur from the gas oil feed varies significantly depending on the 
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pressure at which the unit operates. FCC pretreaters can generally be subdivided into high 

pressure units (1400 psi and above), medium pressure units (900 to 1,400 psi), and low pressure 

units (below 900 psi). High pressure FCC pretreaters are capable of removing about 90 percent 

of the sulfur contained in the gas oil feedstock to the FCC unit, while low and medium pressure 

units are capable of removing 65 to 80 percent of the feed sulfur. Information we received from 

the vendors also indicated that refiners with both a pretreater and a postreater are producing FCC 

naphtha that ranges from 200 to 450 ppm before being processed by the postreater. Having a 

postreater allows these refineries to not have to operate their pretreaters at a high severity as the 

sulfur will further be reduced to levels necessary to meet the applicable standards in the posttreating units. Refiners with only a pretreater are making lower sulfur FCC naphtha in the 75 to 

100 ppm range, according to vendor estimates.

With this information we used our refinery-by-refinery model to estimate the pretreater 

desulfurization rates required to get FCC naphtha sulfur levels within the ranges specified. We 

estimated that FCC units with a pretreater and a naphtha postreater are operating their pretreaters 

at a severity which results in a 76 percent desulfurization of the FCC feed stream. This number 

represents the national average. While the actual severity at which the pretreating units are run 

varies on a refinery-by-refinery basis this average was used in our modeling for all refineries 

with both pretreating and postreating units due to a lack of refinery-specific information. For 

FCC units with a feed pretreater but no postreater we calculated the FCC naphtha sulfur level 

required by refiner to make a refinery gate gasoline that meets the Tier 2 standard. To do this 

calculation we used the gasoline yields from our refinery-by-refinery model along with the 

gasoline blendstock sulfur levels discussed above and shown in Table 5-18. These calculations 

showed that refiners with FCC feed pretreating units, but no postreaters, need to produce FCC 

naphtha that averages about 85 ppm on a national level. This sulfur level corresponded to these 

refiners operating their pretreaters at a severity that results in a reduction of sulfur in the FCC 

feed stream of approximately 91-92 percent. This number is close to the estimate we received 

from the vendors for this category of refineries and therefore was used in our refinery-byrefinery model to determine the FCC feed sulfur level for refiners with pretreaters.

After we have calculated the sulfur level of the FCC feed we must then take into 

consideration the impact the FCC unit itself has on the sulfur level of the FCC naphtha. We 

reviewed several literature sources7,8and found that the FCC naphtha sulfur level can be 

accurately determined by dividing the FCC feed sulfur level by 20 for refineries with an FCC 

feed pretreating unit. For refineries without an FCC feed pretreater, the FCC naphtha sulfur 

levels can be calculated by dividing the desulfurized FCC feed sulfur level by 10. In these cases 

the effect of the FCC unit itself on the sulfur level of the FCC naphtha is lower, as the FCC feed 

has already been through a desulfurization process. These factors, when combined with the 

sulfur levels of the FCC streams as discussed above, allow us to calculate the sulfur level of the 

FCC naphtha before any post-treating operations on a refinery-by-refinery basis. The results of 

this analysis are summarized in Table 5-24 below.
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Table 5-24 FCC Naphtha Sulfur Levels for Various Refinery Configurations

PADD 1 PADD 2 PADD 3 PADD 4/5

No Pretreater or Postreater N.A.

a

-<3 b606 784

Pretreater Only N.A.

a

73 74 82

Postreater Only 1348 1157 705 1036

Pretreater and Postreater 49 <3 b278 230

a N.A. – not applicable, no units of this type in the PADD

b

Since there are less than three refiners in this PADD with the described configuration, the data was 

removed to protect potential CBI concerns.

This information, along with the information described in previous sections (e.g., 

gasoline blendstock volumes and sulfur levels, desulfurization equipment currently in place at 

refineries, and crude oil sulfur levels) allows us to conduct the best analysis for the baseline case 

for our refinery-by-refinery model. This baseline case reflects what we believe the current 

operating conditions are at each refinery, including any modifications they have made to meet 

the Tier 2 sulfur standards. The next step of our analysis was to project what further changes, 

either to operations, adding new equipment or revamping existing units, each refiner would have 

to make to meet the proposed Tier 3 standards. After these changes are estimated, we can then 

estimate the cost associated with each of these changes, and ultimately the cost of the program.

5.1.4 Cost Inputs for the Sulfur Control Technologies

After we determined the sulfur levels of the gasoline blendstocks for each refinery and 

the sulfur levels that these blendstocks would have to achieve to meet the proposed Tier 3 sulfur 

standards, the next step in our refinery-by-refinery analysis was to project the changes to refinery

units and operations each refinery would have to make in order to comply with the proposed Tier 

3 sulfur standard. The first step refiners would take for further gasoline sulfur control would be 

to desulfurize the light straight run, natural gas liquids, and butane blendstocks. The costs to 

reduce the sulfur content of these streams is relatively low and would therefore be the most cost 

effective way to further reduce the sulfur content of the finished gasoline. We also projected in 

our analysis for RFS2 that 50 percent of all gasoline produced by refiners in 2017 would contain 

15 percent ethanol. In 2030 we projected in our analysis for RFS 2 that almost all gasoline 

contains 15 percent ethanol.G Because ethanol tends to be a relatively low sulfur blendstock 

(assumed to be 5 ppm for our refinery-by-refinery model) increasing the amount of ethanol in the 

gasoline pool lowers the overall sulfur content of the gasoline. Reducing the sulfur content of 

the LSR, NGL, and butane streams and increasing the ethanol content, however, would be 

insufficient to allow refiners to comply with the proposed Tier 3 standards. Refineries with an 

FCC unit would still have to reduce the sulfur content of their FCC naphtha blendstock in order 

to meet the Tier 3 standards.

 

G

In our Tier 3 cost analysis, we inadvertently only accounted for the ethanol volumes that are contained in the 2009 

RFG database, not those mandated by RFS 2. For our final rulemaking cost analyiss, we will adjust our model to 

account for higher ethanol use in the future, which will slightly reduce our Tier 3 cost estimates, due to the dilution 

effects of ethanol. 
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For each refinery we considered two cases. In the first case each refinery had to meet the 

proposed Tier 3 gasoline sulfur standard of 10 ppm. In order to meet this standard, as discussed 

in 0, we determined that they would have to reduce the sulfur level of their FCC naphtha stream 

to 25 ppm. We also considered a case where each refinery would reduce the sulfur level of their 

gasoline to 5 ppm. We assumed that to meet a 5-ppm sulfur standard, refiners would desulfurize 

the butane blendstock to 10ppm in addition to further lowering sulfur in the FCC naphtha to 10 

ppm. This information was used to help us determine which refineries might reduce the sulfur 

level of their gasoline below our proposed 10-ppm standard to earn credits for our ABT 

scenarios discussed in Sections 5.2.1.2 and 5.2.1.3.

Our refinery-by-refinery model assumed that reducing the sulfur content of the FCC 

naphtha to 25 ppm and 10 ppm for the two cases discussed above would require that each 

refinery that produces FCC naphtha have an FCC naphtha postreater. In our final rule we plan to 

investigate the ability of refineries with high pressure pretreaters to meet the required standards 

without the addition of postreaters. For companies that already have an FCC naphtha postreater 

we assumed that all that would be necessary to meet the proposed sulfur standards was to revamp 

their existing FCC postreating units. We received cost information from several vendors for 

revamping FCC postreating units and assumed a revamp cost for each refinery in line with the 

cost projections quoted by the vendor of the technology already in place in their refinery. We 

assumed that refineries with FCC units that currently do not have an FCC postreater would have 

no choice other than to add a new grass roots FCC postreating unit. We ultimately only received 

cost information from one vendor for the cost of adding a new grass roots FCC postreating unit 

that fit the sulfur reduction requirements of the proposed Tier 3 program. We therefore assumed 

the cost for each refinery that would need to add a new FCC post-treating unit would be in line 

with this estimate. More details on the costs used in our refinery-by-refinery model for the 

desulfurization of LSR, NGL, and butanes as well as the new FCC post-treating units and 

revamps can be found in the following sections.

5.1.4.1 FCC Naphtha Desulfurization Costs

To estimate the cost for revamping existing FCC postreating units or for adding new 

postreating capacity, we contacted several technology vendors for cost estimates and reviewed 

literature, including cost information provided for the Tier 2 rulemaking. Because no two 

refineries are exactly the same, the cost for new FCC postreater units or revamps to existing units 

will vary significantly from refinery to refinery. Some of the factors that have the most 

significant impact on the cost of FCC postreaters are the technology that the refiner used to 

comply with Tier 2, the volume of FCC naphtha, the sulfur content of the FCC unit feed and the 

level of desulfurization in the existing postreater, and the location of the refinery. Based on 

feedback from vendors we considered three categories of FCC postreaters based on whether the 

FCC naphtha (the feed for the existing Tier 2 postreater) contained low (0 – 400 ppm) medium 

(400 – 1,200 ppm) or high (>1,200 ppm) levels of sulfur. Specific cost factors applicable to 

estimating unit revamps or grass units were also taken into account as described below. The 

following sections discuss in greater detail how the cost estimates we received from the vendors 

were used in our refinery-by-refinery analysis.
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5.1.4.1.1 Cost to Revamp Existing FCC Naphtha Postreaters

We obtained information from several technology providers for the revamp costs of 

existing FCC postreaters. One of the estimates submitted was deemed not credible as inadequate 

information was provided and the capital and operating costs were extraordinarily high relative 

to the rest of the cost information we received. We believe it is likely that this cost estimate 

represented not only a grass roots FCC postreater, but also significant refinery investment in 

other refinery processes such as FCC feed pretreating, coker unit expansion etc. As these costs 

did not seem to be a reasonable representation of the revamp costs for FCC post treating, this 

estimate was not considered in our FCC postreater cost analysis. Some of the submitted 

information only had cost information for our medium sulfur (400 – 1,200 ppm) FCC feed case. 

Another technology provider did not provide cost estimates for producing an FCC naphtha with a 

sulfur level of 25 ppm, corresponding to a finished gasoline with 10 ppm sulfur, therefore, we 

needed to interpolate their cost information. Because the cost information provided by the 

technology providers was labeled CBI, this cost information cannot be listed individually, 

however we aggregated the cost information we received for FCC postreater revamps to meet 10 

ppm and 5 ppm sulfur levels in gasoline. The aggregated information is summarized in Table 

5-26 and Table 5-27.

One of the vendors we contacted for a cost estimate for FCC naphtha desulfurization 

technology provided information for several potential FCC postreater revamp cases. The first 

case was a no capital costs case where refiners made no equipment modifications, but rather 

solely made operational changes using their existing equipment installed for Tier 2. The second 

case we requested was one where refiners would incur only minor capital costs and was intended 

to be used for analyzing program options with moderate octane costs. The third case we 

requested was one where refiners were willing to incur greater capital costs in order to minimize 

operating costs and octane loss. The majority of the vendors only supplied cost estimates for the 

third case, which included adding an additional catalyst reactor bed to the existing FCC 

postreater unit. This ensures that refiners will be able to run their existing FCC postreater at 4 to 

5 year catalyst cycle lengths, which is a critical feature for FCC unit operations.

The costs for the FCC postreater revamps submitted by one of the vendors, however 

showed that for low (0 – 400 ppm) and medium (400 – 1,200 ppm) sulfur FCC naphtha sulfur 

levels, second case, with low capital costs, resulted in the lowest cents per gallon costs for 

meeting the proposed 10-ppm Tier 3 standards. According to this vendor these cases also had a 

4 to 5 year catalyst cycle length, equivalent to the higher capital cost cases even though a second 

stage reactor was not required. We therefore assumed that refineries using this vendor’s 

technology would choose the minor capital cost pathway for meeting the 10-ppm Tier 3 standard 

when they had low or medium sulfur levels in their FCC feed. The high capital cost cases for 

producing gasoline to meet a 5 ppm sulfur standard from low and medium sulfur FCC feeds were 

found to have the lowest cost on a cents per gallon basis and were therefore selected by our 

model for these cases.

One vendor only submitted information for postreater revamp cost estimates for FCC 

naphtha in the 400 – 1,200 ppm sulfur category that produced a 5-ppm sulfur gasoline. In our 

refinery-by-refinery model, however, we had multiple refineries with FCC feed sulfur levels in 

the 0 – 400 and >1,200 ppm categories that use this vendor’s postreating technology. In order to 
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apply this vendor’s cost estimate to cases of low (0 – 400 ppm) and high (>1,200 ppm) sulfur 

feed categories we adjusted this vendors 400 – 1,200 ppm postreater revamp cost based on the 

cost differentials between the three FCC naphtha sulfur levels in the other vendors’ revamp 

estimates. We similarly derived a postreater revamp cost estimate to produce a 10 ppm gasoline 

for this vendor using cost differentials between the 10 and 5 ppm cases from other vendors. For 

refineries currently employing technology by other vendors for which we had no specific cost 

information, we used an average of all of the vendors’ estimates to represent FCC postreater 

revamp costs for refiners using this particular technology in our refinery-by-refiner model.

After we had determined cost estimates for the FCC postreater revamps based on 

information from the vendors the next step was to scale these costs based on the size of the FCC 

postreating unit present in each refinery. The vendor estimates submitted for revamp costs were 

based on various FCC postreater design volumes ranging from 10,000 BPSD to 45,000 BPSD 

depending on the base unit size used by the vendor. To determine how to apply these vendor

costs to each refinery, we first calculated each refinery’s maximum FCC naphtha production. 

The maximum production was derived by assuming each refiner runs its FCC unit at its 

maximum nameplate throughput capacity (barrels per stream day) with the FCC naphtha yield 

rates discussed in Section 5.1.3.1.2. We then increased the size of the FCC postreater by 7.5

percent above each refinery’s maximum FCC gasoline production rate as an over design factor to 

account for excess capacity that refiners generally design into their unit for processing additional 

flows of FCC naphtha (i.e, rerunning high sulfur FCC naphtha batches either from that refinery

or from neighboring refineries). After sizing the FCC postreater that would be required for each 

refinery we then scaled the costs given by the vendors using the six-tenths rule as shown below. 

This is a “rule of thumb” cost estimating tool commonly used for cost estimating by the refining 

and petrochemical industries for estimating the cost of a process unit based on a similar unit of 

differing size. 

Cost to Revamp an FCC Unit= A * (B/C)0.6

Where:

A = Cost Estimate Received from the Vendor

B = Size of the FCC Unit in the Refinery

C = Size of the FCC Unit in the Vendor’s Estimate

Equation 5-4 Six-Tenths Rule for Estimating Capital Cost

We also adjusted the costs submitted by the technology providers based on the location of 

each refinery. We assumed that each vendor’s estimate was based on revamping an FCC 

postreater in PADD 3 (Gulf Coast), which is the lowest cost region for installing new capital in 

refineries. The cost for refineries that are not located in PADD 3 were adjusted upwards based 

on a ratio of the cost of refinery capital projects in the PADD in which they are located relative 

to PADD 3. An additional factor was applied to account for the “offsite” costs that are incurred 

when installing new capital in refineries. When vendors provide a cost estimate for their 

technology, this estimated cost is called the inside battery limits (ISBL) cost and it is solely for 

the immediate unit of interest. However, refiners may need to install peripheral capital to 

support the new unit, such as electrical switchgear, a control room, storage for feed, intermediate 

or unit products, and longer than anticipated pipeline runs – these costs are usually considered 

Outside battery limit (OSBL) costs, or offsite costs. In some cases, OSBL costs may include 
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hydrogen and sulfur plant costs, although, for our analysis, we separately estimated the cost for 

providing additional hydrogen and for processing the removed sulfur and included this cost in 

our cost analysis. We estimate that the grassroots offsiste cost factors for very modest levels of 

desulfurization inherent in this analysis are in the 1.2 to 1.3 range, and those for these very

simple revamps are in the 1.1 to 1.15 range (the ISBL costs are multiplied by these factors to 

derive a total cost). These cost factors, as well as the utility prices that we used in our refineryby-refinery cost model, are shown in Table 5-25 below:

Table 5-25 Cost Factors for Various PADDs

PADD 1 PADD 2 PADD 3 PADD 4 PADD 5a

Capital Cost Factor 1.5 1.3 1.0 1.4 1.2

Natural Gas ($/MMBTU) 7.88 7.86 5.64 7.06 7.44

Electricity (¢/kW-hr) 7.98 5.47 5.78 4.69 8.36

Steam ($/1,000 lb) 12.9 12.6 9.20 11.6 11.9

Offsite Capital Cost Factor –

New Units

1.26 1.26 1.20 1.30 1.30

Offsite Capital Cost Factor –

Unit Revamps

1.13 1.13 1.10 1.15 1.15

a Excluding California

The volume-weighted cost estimates for revamping FCC postreaters across the entire 

refining industry as calculated by our refinery-by-refinery model are shown in Table 5-26 and 

Table 5-27 below. These costs are aggregated cost estimates for the FCC revamp costs used in 

our refinery-by-refinery model. In our model, we paired vendor cost data with refineries that are 

already using that particular vendor’s technology for their FCC postreating units. We further 

tailored the information provided by the vendors to match the specific refinery configuration to 

the extent possible. The information that we received from the vendors and the individual 

refinery capital costs, however, cannot be shown due to CBI concerns.

Table 5-26 Revamp Cost for a 30,000 BPSD FCC Postreater for 10-ppm Standard

FCC Feed Sulfur Levela0 – 400 ppm 400 – 1,200 ppm >1,200 ppm Volume Weighted 

Averageb

Capital Cost ($/B ISBL) 235 382 505 265

Hydrogen (scf/bbl) 36.8 25.9 25.9 31.3

Fuel Gas (kBTU/bbl) 15.5 15.4 8.51 14.4

Electricity (kWh/bbl) 0.14 0.27 0.52 0.223

Octane Loss (R+M)/2 0.56 0.46 0.42 0.492

Olefin Decrease (vol%) 2.67 2.14 1.97 2.39

Catalyst Cost ($/B) 0.01 0.01 0.01 0.010

Steam (lb/bbl) 0.61 30.88 0.61 13.44

a $/B = dollars per barrel, scf/bbl = standard cubic feet per barrel; kBTU/bbl = thousand BTU per barrel; kWh/bbl = 

kilowatt-hours per barrel; (R+M)/2 = (research octane + motor octane)/2; vol% = volume percent; $/B = dollars per 

barrel; lb/bbl = pounds of steam per barrel of feed.

b

 Of the refineries that are expected to revamp their FCC naphtha hydrotreater for the no ABT case, 36 have FCC 

naphtha sulfur levels in the 0 – 400 ppm range, 20 have FCC naphtha sulfur levels in the 400 – 1200 ppm range and 

13 have FCC naphtha sulfur levels greater than 1200 ppm. 
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Table 5-27 Revamp Cost for a 30,000 BPSD FCC Postreater for 5 ppm Standarda

FCC Feed Sulfur Level 0 – 400 ppm 400 – 1,200 ppm >1,200 ppm Volume Weighted 

Average

Capital Cost ($/B ISBL) 400 578 600 444

Hydrogen (scf/bbl) 36.2 52.6 56.3 38

Fuel Gas (KBTU/bbl) 22.7 14.6 13.2 17.2

Electricity (kWh/bbl) 0.30 0.52 0.52 0.42?

Octane Loss (R+M)/2 0.60 0.77 0.90 0.62

Olefin Decrease 2.81 3.38 3.96 2.37

Catalyst Cost ($/B) 0.00 0.01 0.00 0.004

Steam (lb/bbl) 23.61 30.88 0.61 26.9

a Assumes every refinery is complying with a 5 ppm gasoline sulfur standard

We also found that there were 11 refineries that had an existing FCC postreaters that 

were not sized large enough to process their maximum FCC naphtha production volume. For 

these refineries we assigned additional capital costs to debottleneck the existing first stage 

reactor in order to increase the postreater capacity so that it could accommodate maximum FCC 

naphtha production. For each refinery with an existing unit that could not process more than 70

percent of our estimate of a refiner’s maximum FCC naphtha production we added capital costs 

to revamp and expand the first stage to increase its capacity to allow the postreater to process 

100 percent of its maximum FCC naphtha rate. For the capital costs for this debottlenecking we 

used 35 percent of the cost of a new grass roots unit (discussed below) for the volume of the 

expansion. We once again used the six-tenths rule to adjust the capital cost for the volume 

expansion needed versus the cost for the 30,000 BPSD grass roots treater used for technology

vendor estimates.

5.1.4.1.2 Cost for Grassroots FCC Postreaters

While all refineries that already have FCC postreaters should be able to meet the 

proposed Tier 3 standards by revamping their existing postreaters, refineries that do not currently 

have an FCC postreater would have to add a grass roots FCC postreater. To determine the cost 

of building grass roots FCC postreating units at a refinery we similarly requested cost estimates 

from vendors. Only one of the vendors that supplied FCC postreating equipment submitted 

information on the cost of a grass roots FCC postreating unit for desulfurizing FCC naphtha with 

a feed sulfur content of 100 ppm. Based on the calculation methodology shown in Equation 5-2, 

we estimate that refineries that require a grass roots postreater will already have an FCC feed 

sulfur level that averages between 85 and 100 ppm as these refineries already have FCC feed 

pretreaters. The other grassroots vendor estimate we received, as well as those we received for 

Tier 2, represented a grass roots postreater with an FCC feed sulfur content of about 800 ppm. 

These estimates were deemed to be not representative of the costs to refineries that would be 

installing grass roots postreating units as the capital, hydrogen, and other operating costs would 

be much higher for an FCC feed sulfur of 800 ppm vs. 100 ppm. We did not consider this other 

vendor’s cost estimate for a grass roots postreater and therefore relies on a single vendor’s cost 

estimate for grass roots FCC postreating units for the Tier 3 program. In our FRM analysis we 

attempt to incorporate additional information from other vendors and literature sources. This one 
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vendor cost estimate seemed reasonable relative to the other cost data that we have for higher 

levels of desulfurization.

The vendor estimate submitted for a grass roots postreater was based on a postreater with 

a capacity of 30,000 BPSD capable of producing an FCC naphtha with a sulfur level of 10 ppm, 

corresponding to a gasoline sulfur level of 5 ppm. To scale the cost submitted by the technology

vendor to be applicable to a specific refinery, we used a similar methodology to that which was 

used for postreater revamps. We first determined the appropriate size for each unit based on 

each refiners maximum FCC naphtha production rate, adding 7.5 percent above each refinery’s 

maximum FCC gasoline production rate as an over design factor. We then used the six-tenths 

rule (Equation 5-4) to scale the cost reported by the vendor up or down as appropriate based on 

the relative volume of the grass roots unit required by the refinery and the size on which the 

vendor’s cost estimate was based. We once again assumed that the capital cost from the 

technology vendor was representative of a refinery in PADD 3 and adjusted the cost based on the 

cost of refinery capital projects in the PADD in which they are located relative to PADD 3. 

Finally, we used a new unit offsite adjustment factor as listed in Table 5-25 to determine the final 

cost of a grass roots FCC postreater for each refinery. The costs to produce an FCC naphtha 

with a sulfur level of 25 ppm (corresponding to a 10-ppm gasoline) were estimated based on the 

grass roots postreater unit that makes FCC naphtha for the 5-ppm standard. These costs are 

summarized in Table 5-28 and Table 5-29 below.

Table 5-28 Cost for a 30,000 BPSD Grass Roots FCC Postreater for 10-ppm Standard

FCC Feed Sulfur Level 100 ppm

Capital Cost ($/B ISBL) 1500

Hydrogen (scf/bbl) 62.1

Fuel Gas (KBTU/bbl) 4.95

Electricity (kWh/bbl) 1.06

Octane Loss (R+M)/2 0.55

Olefin Decrease 2.65

Catalyst Cost ($/B) 0.04

Steam (lb/bbl) 20.0
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Table 5-29 Cost for a 30,000 BPSD Grass Roots FCC Postreater for 5-ppm Standard

FCC Feed Sulfur Level 100 ppm

Capital Cost ($/B ISBL) 1500

Hydrogen (scf/bbl) 133

Fuel Gas (kBTU/bbl) 9

Electricity (kWh/bbl) 1.06

Octane Loss (R+M)/2 1.00

Olefin Decrease 5.15

Catalyst Cost ($/B) 0.04

Steam (lb/bbl) 20.0

5.1.4.2 Light Straight Run and Natural Gas Liquids Desulfurization Costs

Another action refiners may need to take to reduce the sulfur content of their gasoline is 

to desulfurize their light straight run naphtha (LSR) and natural gas liquids (NGL) blendstocks. 

While these blendstocks have lower sulfur contents than the FCC naphtha, they may be cheaper 

to desulfurize for refineries that are not already treating these streams. Many refineries can 

desulfurize some or all of these blendstocks using existing excess capacity in their naphtha 

hydrotreaters or FCC naphtha postreaters. Further, as opposed to hydrotreating FCC naphtha

which contains olefins, the LSR and NGL blendstocks contain no olefins and therefore, 

hydrotreating them does not result in octane loss and has a lower hydrogen consumption. The 

combination of the potential for using excess capacity in existing units and low operating costs 

result in the relatively low desulfurization costs for the LSR and NGL blendstocks. From our 

discussions with refiners, several refineries indicated that they would install new standalone 

hydrotreaters for processing LSR and NGL blendstocks, though it is unclear which other 

refineries will have to add equipment to desulfurize LSR and NGL. For our Tier 3 cost 

estimates, however, we have conservatively used our refinery-by-refinery model to estimate the 

costs for other refiners to hydrotreat all refiner volumes of LSR and NGL by installing additional 

hydrotreating equipment. To determine the cost to desulfurize the LSR and NGL blendstocks we 

first had to determine the volume of blendstock that requires desulfurization. Our determination 

of the quantity of LSR and NGL used as gasoline blendstock at each refinery is discussed in 

Section 5.1.3.1.2. From this total we then subtracted the volume of LSR processed in the 

isomerization unit. Because the isomerization units require a low sulfur feedstock we assume all 

the feed to this unit is treated by the naphtha hydrotreater.

The next step in our assessment of the desulfurization costs of the LSR and NGL 

blendstocks was to estimate the extent to which these blendstocks are already being treated at 

refineries to meet the existing Tier 2 sulfur standards. Based on our discussion with refining 

consultants, vendors and refiners, it appears that in response to the Tier 2 standards refiners have 

altered their operations to use excess capacity in their FCC naphtha postreaters and naphtha

hydrotreaters to reduce the sulfur content of LSR and NGL blendstocks. Since information on 

the extent to which these streams are currently being hydrotreated is not publicly available we 

estimated these volumes using the capacities of the FCC postreaters and reformer feed 

hydrotreaters under normal refiner crude throughputs and yields from the refinery-by-refinery

model.
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We evaluated each refinery’s capacity to hydrotreat LSR and NGL using existing 

equipment by first determining the volume that can be processed in their naphtha hydrotreaters. 

We assumed that every refinery’s production volume of straight run and coker naphthas, as 

determined in Section 5.1.3.1, are processed in the refinery’s reformer feed hydrotreater and that 

any capacity up to 85 percent of the unit’s maximum capacity can be used to treat LSR or NGL. 

If a refinery had insufficient excess capacity in their naphtha hydrotreater to treat all of the LSR 

and NGL volumes we next determined if there was excess capacity in that refinery’s FCC 

postreater. We allowed LSR and NGL to be processed using excess FCC postreating capacity in 

refineries where the capacity of the FCC postreater exceeds 120 percent of that required to 

process a refinery’s maximum FCC naphtha yield as determined by the refinery-by-refinery

model. Only several refiners had excess FCC postreating capacity available for the treating of 

LSR or NGL feedstocks, as the capacity of most FCC postreaters was less than 120 percent of 

the maximum FCC naphtha production rate. We assumed that refiners are currently using any 

excess hydrotreating capacity in their naphtha hydrotreating and FCC postreating units to 

desulfurize LSR and NGL in response to the Tier 2 sulfur standards. To meet the proposed Tier 

3 sulfur standard 10 ppm and the 5 ppm level we considered for the ABT cases, we assumed that 

refiners would desulfurize all of the LSR and NGL blendstocks. We assumed that LSR and 

NGL purchased by refiners would be hydrotreated in addition to that processed by the 

isomerization unit and that blended directly into the gasoline. To determine which of the 

refineries in our refinery-by-refinery model would need to add hydrotreating capacity, we 

evaluated each refinery based on their maximum crude processing rate and yields. We used our 

model to increase each refinery’s yield of coker naphtha, reformer naphtha, isomerate, LSR and 

NGL to correspond to a crude utilization rate of 100 percent. We then used these maximum 

production volumes to evaluate whether or not each refinery could process all of the LSR and 

NGL using excess naphtha hydrotreater and FCC postreating capacity. If a refinery did not have 

sufficient excess hydrotreating capacity for all of the LSR and NGL in these units we assumed 

the refinery would have to either revamp their existing equipment or add new hydrotreating 

capacity. If the additional capacity needed at any given refinery exceeded the existing naphtha

hydrotreater capacity by less than 30 percent we assumed the necessary capacity could be added 

by revamping the existing unit. (For our FRM, we will also evaluate if refiners could also use the 

caustic Merox extraction process to lower the sulfur level in LSR and NGL blendstocks, and if 

so, are they already doing that today). If, however, the additional capacity required exceeded the 

existing reformer feed hydrotreater capacity by more than 30 percent we assumed the refinery

would install a new stand-alone hydrotreater to desulfurize the excess LSR and NGL. Based on 

available capacity in our refinery-by-refinery model and 2009 crude throughput data, we 

estimated that refiners are already hydrotreating 66 percent of the LSR and NGL that are directly 

blended into gasoline (excluding LSR processed in the isomerization units). The results of this 

assessment are shown in Table 5-30 below.

Table 5-30 Refineries Adding Hydrotreating Capacity for LSR and NGL

New Hydrotreater 

(No FCC Unit)

New Hydrotreater 

(with FCC Unit)

Revamped 

Hydrotreater 

(No FCC Unit)

Revamped 

Hydrotreater 

(with FCC Unit)

Number of 

Refineries

4 11 1 7
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We conservatively evaluated the capital cost required for hydrotreater revamps and new 

units by assuming that refiners will size their hydrotreater equipment needs to treat all production 

volumes of LSR and NGL based on each refinery’s maximum crude run rate. The operating 

costs used in our refinery-by-refinery model, however, are based on LSR and NGL blendstock 

rates from the models yields at the 2009 operational crude throughputs as discussed in Section 

5.1.3.1. Sizing the equipment this way allows each refiner to have excess hydrotreater capacity 

utilization, which is beneficial in the event of process unit shutdowns and to reprocess 

blendstocks from abnormal operations.

Our estimate for the cost of adding a new hydrotreater at a refinery was obtained from 

Gary and Handework’s Petroleum Refining Technology and Economics, page 182-183, Curve C, 

Table 9.1, 30,000 BSD unit. The capital cost for a grass roots hydrotreater listed by this source 

was for a hydrotreater with a capacity of 30,000 BPSD and was based on 1999 dollars. We 

multiplied this cost by 1.534 to determine the equivalent cost in 2010 dollars based on the 

relative increase in the Nelson Refining Construction index from 1999 to 2010 (listed as 1497 

and 2296 respectively). We used the six-tenths rule to scale the capital cost listed in Petroleum 

Refining Technology and Economics to those of differing capacities based on relative size of the 

desired unit. We assumed a hydrogen consumption of 40 SCF/Bbl for the processing of LSR 

and NGL blendstocks which we obtained from the Jacobs Refining LP modeling database for 

naphtha hydrotreating as this information was not presented in the literature source. For 

refineries that only required a revamp of existing units we assumed a capital cost equivalent to 

40 percent of the cost of a new hydrotreating unit of equal size. We assumed equivalent 

operating costs for new hydrotreating units and revamped units. The capital and operating costs 

for these hydrotreating units that were used in our model is shown in Table 5-31 below.

Table 5-31 Capital and Operating Costs for LSR and NGL Hydrotreaters

New Hydrotreating Units Revamped Hydrotreating Units

Capital ($/BBL ISBL) 1380 550

Hydrogen (SCF/BBL) 40 40

Fuel Gas (kBTU/BBL) 100 100

Electricity (kWh/BBL) 2.0 2.0

Octane Loss (R+M)/2 0.0 0.0

Change in Olefins (vol%) 0.0 0.0

Steam (lb/BBL) 6.0 6.0

Catalyst Cost ($/BBL) 0.03 0.03

5.1.4.3 Butane Desulfurization Costs

After considering the desulfurization of LSR and NGL we then estimated the cost to each 

refinery for desulfurizing the butane that is blended into gasoline. For the baseline case without 

ABT where each refinery had to meet the 10-ppm sulfur standard we did not assume any cost 

associated with butane desulfurization, as we assumed that any refiners likely to reduce the 

sulfur content of the butane stream already had the necessary equipment. To estimate the costs 

for each refinery to produce a 5-ppm gasoline for our ABT cases, however, we assumed 

refineries would reduce the sulfur content of any butane blended into gasoline using extractive 

desulfurization units (such as Merox or Merichem) which extract mercaptan sulfur (the bulk of 
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sulfur species in butanes) from butane volumes that are added to each refiner’s gasoline pool. 

We accounted for butane treating cost in our refinery-by-refinery model by adding capital to 

install Merox treating units to each refinery. This is likely a conservative estimate based on our 

lack of information on which refineries may already have these units. We set up the model to 

size the Merox treater based on each refinery’s maximum addition of butane volume to the 

gasoline pool, using the amount that is blended in the winter, as winter grade gasoline has the 

highest seasonal gasoline RVP limits and therefore allows for the highest levels of butane 

blending. As with some of the LSR and NGL desulfurization units, it is unclear if additional 

butane desulfurization units will be added to meet the Tier 3 standards. As a conservative 

estimate, though, we have included the cost of adding these units to our program costs.

To calculate the volume of butanes added to winter grade gasoline we used the refineryby-refinery model estimate of the full year addition of butane to each refiners CG and RFG 

gasoline grades which was based on the EPA annual gasoline database. The winter addition of 

butane volumes was assumed to be 66 percent of the total annual volume of butanes added to a 

refiner’s gasoline pool on a yearly basis. The capital cost we used was for a 10,000 BPSD 

Merox unit and was listed in 1995 dollars.

9

 We multiplied this cost by 1.64 to determine the 

equivalent cost in 2010 dollars based on the relative increase in the Nelson Refining 

Construction index from 1995 to 2010. We then multiplied the maximum winter daily butane 

volume by 1.08 to account for an over design factor to size the equipment. We next used the sixtenths rule to scale the capital cost listed in Petroleum Refining Technology and Economics to 

those of differing capacities based on relative size of the desired unit in BPSD divided by 10,000 

(the size of the unit the cost is based on). Finally, we applied offsite and PADD location factor 

adjustments as listed in Table 5-25. For the operating cost we assumed that each unit had an 

annual fixed operating cost of 6.7 percent of the total installed capital cost. To this we added an 

operating cost of 0.03¢ per gallon based on an operating cost figure from literature listed in 1995 

dollars and scaled to 2010 dollars using appropriate factors. Operating costs were presumed to 

represent all process energy and utility costs, as well as costs associated with purchasing caustic. 

Table 5-32 Merox Treating cost for Sulfur Removal from Butane

Size, BPSD Capital costs, $ MM Operating Costs, ¢/gal

10,000 3.28 0.03

5.1.5 Overall Cost Methodology

5.1.5.1 Sulfur Costs without ABT Program

While we are proposing an ABT program as described in Section V.E.1 of the preamble, 

we also evaluated a scenario with no averaging, banking, or trading between refineries. Under 

this scenario, every individual refinery would comply with the 10-ppm annual average sulfur 

standard, and none would lower their annual average sulfur below 10 ppm. This evaluation 

provided a reference point for determining the impacts of an ABT program on compliance costs. 

In complying with the 10-ppm sulfur standard, our refinery-by-refinery analysis 

concluded that some refineries would do so with revamps to existing equipment while others 

would need to install new grassroots equipment. Those refineries whose average sulfur levels 

are already at or below the 10-ppm standard would not need to do anything. Because the most 
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recently available data was for 2009 when some phase-ins under Tier 2 were still effective, we 

made adjustments to ensure that refinery-specific sulfur levels represented levels after all Tier 2 

phase-ins had expired. For instance, if any refinery exhibited an annual average sulfur level 

above the 80-ppm cap based on 2009 data, we assumed that their current annual average sulfur 

level was 30 ppm. Further discussion of our refinery-by-refinery analysis can be found in 

Section 5.1.3.

To determine the impacts on cost of the 10-ppm sulfur standard without ABT, we 

volume-weighted the refinery-specific ¢/gal costs calculated according to the refinery-byrefinery modeling using 2009 gasoline production data for each refinery. Likewise, we 

determined nationwide capital costs by summing the individual revamp and/or grassroots 

equipment costs across all refineries. In these calculations, we assumed that all early credit 

generation and phase-ins had expired, consistent with our approach to the ABT analysis. The 

results of our analysis of costs under a scenario without ABT are provided in Section 5.2.1.1

below.

5.1.5.2 Sulfur Costs with ABT Program

As described in Section V.E.1 of the preamble, we are proposing an ABT program that is 

designed to ease the overall burden on the industry while still achieving the 10-ppm annual 

average sulfur standard for the nation as a whole. Under the proposed ABT program, refineries 

that can reduce sulfur below 10 ppm at a relatively low cost can generate credits which can then 

be sold to refineries for whom the cost of attaining the 10 ppm sulfur standard would be 

relatively high. The net effect of this credit trading would be expected to reduce the overall cost 

of the program.

To estimate the impact that the ABT program could have on nationwide average fuel 

costs, we began with the refinery-by-refinery costs described in Section 5.2.1.1 for sulfur 

reductions down to either 10 ppm or 5 ppm. We then determined the lowest cost option among 

three alternatives for each refinery:

1. The refinery reduces its sulfur to 10 ppm.H

2. The refinery reduces its sulfur to 5 ppm and generates credits (in ppm-gal) for the 

increment between 10 ppm and 5 ppm.

3. The refinery does not lower sulfur, but instead obtains credits to comply with the 

10-ppm standard.

A fourth category applied to refineries whose average gasoline sulfur levels are already below 10 

ppm. All such refineries were assumed to generate credits for the increment between 10 ppm 

and their current sulfur level.

 

H No refinery compliance margins were included in this analysis.
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To optimize the nationwide average costs under an ABT program, we determined the 

credit price at which the total number of credits generated was equal to the total number of 

credits consumed. To simplify the modeling of how an ABT program might operate, we focused 

on the circumstances that refineries would face in the longer term, specifically after 2020. This 

approach meant that the ABT program modeling did not consider the impact on gasoline sulfur 

levels of delayed compliance for small refiners and small volume refineries, nor did it consider 

the generation and use of any early sulfur credits. Moreover, our ABT modeling considered only 

gasoline sold for use outside of California, and only gasoline produced by domestic refineries 

(not imports). 

Since the cost information available for this NPRM was limited to the costs of reducing 

sulfur to either 10 ppm or 5 ppm (under which FCC gasoline sulfur levels would be reduced to 

either 25 or 10 ppm, respectively), we were not able to estimate refinery-specific costs of 

reducing sulfur to other levels. As a result, our ABT modeling could not account for scenarios in 

which a refinery makes some capital investments to lower sulfur to some interim level, such as 

20 ppm, and then obtains credits in order to demonstrate compliance with the 10-ppm standard. 

Our ABT analysis also could not account for credit generation at sulfur levels other than 5 ppm. 

Our ABT analysis, then, most likely underestimates the cost savings that could occur due to ABT 

since the greatest efficiencies are achieved when every refinery has the option of using any 

combination of capital investments and credits generation or use. For the final rule, we may

investigate methods for expanding our ABT analysis to examine these types of scenarios.

We evaluated two ABT scenarios designed to bound the likely outcomes. In the first 

scenario, we assumed perfect nationwide credit trading in which all credit generators make their 

credits available to any refinery that needs them, regardless of their ownership. As described in 

Section V.E.1 of the preamble, the proposed ABT program is designed such that it could operate 

in this way. However, such perfect nationwide trading may not be realistic. Under Tier 2 today, 

there is still a considerable amount of inter-company trading occurring, but a significant fraction 

of Tier 2 sulfur credits are bought and sold within companies.I Under Tier 3, it will be more 

difficult to generate credits, and also more difficult to make up for deficits. Consequently, we 

also investigated a second ABT scenario in which trading between companies does not occur and 

averaging would only occur within companies that own more than one refinery. Under this 

second scenario, individual companies might decide to bank credits for their own use, declining 

to make credits that they generate at one of their facilities available to other companies. They

might do this, for instance, to address unplanned equipment downtime or other circumstances 

that could make future compliance more difficult. While we do anticipate some trading to occur 

between companies, we believe that this second, more limited scenario more closely represents 

how our proposed ABT program might actually operate. As a result, the within-company credit 

trading scenario is the primary scenario that we used to estimate overall costs for our proposed 

program. It represents a somewhat conservative scenario with respect to estimation of costs. 

The results of both scenarios are described in Sections 5.2.1.2 and 5.2.1.3 below.

 

I Based on Tier 2 sulfur compliance data, of the 26 companies that purchased sulfur credits in 2010, eight bought 

credits only from their own company, and another five bought credits both from their own company and from other 

companies. The remainder bought credits only from other companies.
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Since an ABT program would allow some refineries to continue producing gasoline with 

an average sulfur level above 10 ppm, we also investigated whether any areas of the country 

might experience inordinately high sulfur levels as a result, and if so whether those higher sulfur 

levels might be problematic in terms of either vehicle performance or local air quality. While we 

were not able to model the distribution of all gasoline from the point of production to the point of 

consumption, we did compare average sulfur levels at each refinery's location to the location of 

nearby ozone nonattainment areas (based on the designation status with respect to the 1997 

ozone NAAQS). A discussion of our analysis of sulfur levels by refinery and location can be 

found in Section 5.2.1 below.

5.2 Estimated Tier 3 Sulfur Control Costs

5.2.1 Sulfur Cost Results

5.2.1.1 Cost of the Sulfur Program without ABT

Without ABT, all refineries would be required to meet the 10-ppm annual average sulfur 

standard. We estimate that 13 refineries are already at or below 10 ppm, and thus their costs of 

compliance would be zero. The remaining refineries would incur ¢/gal costs at various levels 

depending on their particular configurations and the steps they would need to take. While we 

estimate that the nationwide average cost of compliance would be 0.97 ¢/gal under this scenario, 

two refineries would incur costs above 6 ¢/gal. The distribution of costs is shown in Figure 5-1

as a function of the number of refineries.

0

5

10

15

20

25

30

35

40

45

50

0.00 0.50 1.00 1.50 2.00 2.50 3.00 3.50 4.00 4.50 5.00 5.50 6.00 6.50 7.00 7.50

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cents/gallon

Nationwide average = 0.97 c/gal

Figure 5-1 Distribution of Compliance Costs by Refinery for 10 ppm without ABT 

Program
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In order to enable evaluation of the costs of the program with ABT, we also needed to 

estimate the costs for each refiner to reach 5 ppm. The distribution of costs for this scenario is 

shown below.

0

5

10

15

20

25

30

35

40

0.00 0.50 1.00 1.50 2.00 2.50 3.00 3.50 4.00 4.50 5.00 5.50 6.00 6.50 7.00 7.50

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cents/gallon

Nationwide average = 1.38 c/gal

Figure 5-2 Distribution of Compliance Costs by Refinery for 5 ppm without ABT Program

5.2.1.2 Cost of Sulfur Control Program with ABT

Under a scenario in which credit trading only occurs between refineries within the same 

company, we determined that about 23 percent of refineries would be involved in either 

generating or consuming credits. Table 5-33 summarizes our estimate of the amount of credit 

trading that would occur under this scenario.

Table 5-33 Impacts of ABT Program with Intra-Company Credit Transfers

Number of refineries whose sulfur is already below 10 

ppm, and which generate creditsa 3

Number of refineries which lower their sulfur to 5 ppm 

and generate credits 15

Number of refineries that do not lower their sulfur but 

instead consume credits 8

Number of refineries that lower sulfur to 10 ppm and 

neither generate nor consume credits 85

a Additional refineries also have sulfur levels below 10 ppm, but do not generate credits under this scenario.
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The nationwide average cost of compliance with the 10-ppm sulfur standard would be 

reduced from 0.97 ¢/gal without ABT to 0.89 ¢/gal under company-only trading.J However, 

there would continue to be a significant variation in costs by refinery as illustrated in Figure 5-3.

0

5

10

15

20

25

30

35

40

45

50

-0.5 0 0.5 1 1.5 2 2.5 3 3.5 4 4.5 5 5.5 6 6.5

Num

b

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cents/gallon

Nationwide average = 0.89 c/gal

Figure 5-3 Distribution of Compliance Costs by Refinery for 10-ppm Average Standard 

with ABT Program and Intra-Company Credit Transfers

In this figure, refineries with negative costs are those whose sulfur levels were already

below 10 ppm and who would generate and sell credits. While there are several refineries with 

costs above 3 ¢/gal in this scenario, in practice we would expect such refineries to incur lower 

costs by pursuing some combination of lowering their gasoline sulfur to some level above 10 

ppm and also purchasing credits to cover the remaining sulfur reductions needed to reach 10 

ppm. The analysis we conducted for this draft RIA was unable to capture such nuances, but we 

will be investigating them for the final rule.

Figure 5-4 illustrates how the cumulative distribution of costs by volume would change 

under this scenario. In general, the distribution would shift downward slightly in comparison to 

a scenario in which there was a 10 ppm average standard but no ABT program.

 

J Approximately 55% of this desululfurization cost is comprised of the variable cost, and the majority of that is due 

to the cost of recovering the octane lost when the hydrotreater unit saturates the octane-rich olefins contained in the 

FCC naphtha. Another 30% of the cost is due to the capital cost amortized over the volume of gasoline. Finally, 

about 15% of the cost is due to the fixed operating cost which includes the maintenance of the new equipment and 

taxes. 
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0

1

2

3

4

5

6

0 20000 40000 60000 80000 100000 120000

Cents/gallon

Cumulative gasoline volume (mill gallons)

Distribution with no ABT

Distribution with company-only trading

Figure 5-4 Distribution of Costs by Volume for 10-ppm Average Standard with ABT 

Program and Intra-Company Credit Transfers

Since we estimate that 85 out of 111 refineries would meet the 10-ppm sulfur standard 

without the use of credits under the intra-company trading scenario, only about 5 percent of 

gasoline would continue to have an annual average sulfur level above 10 ppm. These refineries 

tend to be smaller than average. The distribution of sulfur levels under this scenario is shown in 

Figure 5-5.
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0

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5 10 15 20 25 30 35 40 45 50 55 60 65 70 75 80 85

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Sulfur (ppm)

Figure 5-5 Distribution of Sulfur Levels by Refinery with ABT Program and IntraCompany Credit Transfers

Of the 5 percent of gasoline volume that would have average sulfur levels above 10 ppm under 

this scenario, nearly half would be below 20 ppm, and a majority would be below 30 ppm. 

Moreover, as noted above, we would expect refineries with the highest sulfur levels to pursue 

some combination of lowering their gasoline sulfur to some level above 10 ppm and also 

purchasing credits to cover the remaining sulfur reductions needed to attain the 10-ppm standard.

We also investigated whether the existence of refineries producing gasoline with average 

sulfur levels higher than 10 ppm after 2020 might contribute to local areas where the average 

sulfur level is higher than 10 ppm, and whether this might be problematic in terms of either 

vehicle performance/emissions or local air quality. To do this, we compared the location of 

refineries with sulfur levels higher than 10 ppm to the location of other refineries serving similar 

areas. We also compared the location of refineries to areas that have historically been in 

nonattainment for ozone.

As shown above in Table 5-33, we estimated that eight refineries would have annual 

average sulfur levels above 10 ppm under company-only credit trading. Three of these refineries 

serve areas that are, at a minimum, several hundred miles away from historical ozone 

nonattainment areas. The remaining five refineries with average sulfur levels above 10 ppm 

supply areas that are also supplied by other refineries whose average sulfur levels would be 10 

ppm or lower. As a result, gasoline with higher sulfur levels would likely be diluted by gasoline 

with lower sulfur levels, and we would not expect any ozone non-attainment area to receive 

gasoline with an average sulfur level higher than 20 ppm. 
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5.2.1.3 Cost of Sulfur Control Program with ABT and Nationwide Credit 

Transfers

Although we believe that credit trading will occur primarily between refineries within the 

same company, we also investigated a scenario in which credit transfers would occur between all 

refineries regardless of location or company. For this scenario, we determined that more than 60

percent of refineries would be involved in either generating or consuming credits. Table 5-34

summarizes our estimate of the amount of credit trading that would occur under this scenario.

Table 5-34 Impacts of ABT Program with Nationwide Credit Transfers

Number of refineries whose sulfur is already below 10 

ppm, and which generate credits 12

Number of refineries which lower their sulfur to 5 ppm 

and generate credits 34

Number of refineries that do not lower their sulfur but 

instead consume credits 25

Number of refineries that lower sulfur to 10 ppm and 

neither generate nor consume credits 40

The nationwide average cost of compliance with the 10-ppm sulfur standard would be reduced 

from 0.97 ¢/gal without any ABT program to 0.79 ¢/gal under nationwide trading. There would 

continue to be some variation in costs by refinery as illustrated in Figure 5-6, though the 

distribution would not be as wide as under the intra-company credit transfer scenario.

0

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-0.5 0.0 0.5 1.0 1.5 2.0 2.5 3.0 3.5 4.0 4.5 5.0 5.5 6.0 6.5

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cents/gallon

Nationwide average = 0.79 c/gal

Figure 5-6 Distribution of Refinery Compliance Costs with ABT Program and Nationwide 

Credit Transfers
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Figure 5-7 illustrates how the cumulative distribution of costs by volume would change 

under this scenario. In comparison to a scenario wherein credit trading occurs only between 

refineries within the same company, the downward shift in the distribution would be more 

pronounced.

0

1

2

3

4

5

6

0 20000 40000 60000 80000 100000 120000

Cents/gallon

Cumulative gasoline volume (mill gallons)

Distribution with no ABT

Distribution with nationwide trading

Figure 5-7 Distribution of Compliance Costs by Volume with ABT Program and 

Nationwide Credit Transfers

Under a scenario in which credit transfers occur nationwide, we estimate that 25 out of 

111 refineries, or about 23 percent, would not lower their sulfur levels and would instead obtain

credits. These refineries tend to be smaller than average, and represent about 17 percent of 

gasoline volume. The distribution of sulfur levels under this scenario is shown in Figure 5-8.
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0

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5 10 15 20 25 30 35 40 45 50 55 60 65 70 75 80 85

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Sulfur (ppm)

Figure 5-8 Distribution of Sulfur Levels by Refinery with ABT Program and Nationwide 

Credit Transfers

Of the 17 percent of gasoline volume that would have average sulfur levels above 10 ppm 

under this scenario, the vast majority would be below 30 ppm. 

As for the previous scenario under which credit transfers would occur only within 

companies, we investigated the localized impact of higher sulfur levels under a scenario in which 

credit transfers occur nationwide and thus there are more refineries with average sulfur levels 

above 10 ppm. Under this scenario, most refineries with average sulfur levels higher than 10 

ppm supply areas that are also supplied by refineries with sulfur levels at 10 ppm or below. We 

expect the dilution of higher sulfur with lower sulfur in such areas to result in annual average 

sulfur levels no higher than about 20 ppm. Of those refineries projected to have average sulfur 

levels higher than 10 ppm under this scenario and which are the primary or only suppliers of 

gasoline to a particular area, none of the affected areas have historically been in nonattainment 

for ozone. 

5.3 Other Cost Studies

Other other cost studies were recently conducted to estimate the cost of additional 

reduction in gasoline sulfur. We evaluated each of these studies and compare them to our own 

cost analysis. 

The International Council for Clean Transportation (ICCT) retained Mathpro in October 

2011 to study the cost of a 10 ppm average gasoline sulfur standard as well as a 1 psi reduction 

in RVP.10 Since the lower RVP standard was modeled as a separate step from the low sulfur 

standard, we were able to isolate the gasoline sulfur reduction costs from the low RVP costs.
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ICCT’s estimated cost for a 10 ppm average gasoline sulfur standard is 0.8 cents per 

gallon which reflects the capital costs amortized assuming a before-tax 7 percent rate of return 

on investment. This cost reflects an assumption that the capital cost for revamps of FCC 

postreaters is 30 percent of the capital costs for a grassroots FCC postreater . Mathpro also 

analyzed costs assuming that the capital cost for revamps of FCC postreaters are 50 percent of a 

grassroots FCC postreaters, which is 1.1 cents per gallon. ICCT’s cost estimate for complying 

with a 10 ppm average gasoline sulfur standard is very close to ours. 

In 2008 The Alliance retained Mathpro to use its LP refinery cost model to estimate the 

costs of what they termed National Clean Gasoline (NCG) in PADDs 1, 2 and 3 (generally 

speaking, this is the part of the U.S. east of the Rocky Mountains).11 Achieving NCG would 

entail reducing gasoline sulfur to 5 ppm under a 10-ppm cap standard and the reduction of 

gasoline RVP to 7 psi. For the low-RVP standard, a 1-psi waiver was allowed for conventional 

gasoline, but not for current low-RVP areas. The study also evaluated two sensitivity cases 

which increases the stringency of the distillation index (DI) from 1250 to 1200. The Alliance 

study also evaluated crude oil price as a second sensitivity case, evaluating crude oil prices at 

$51 /bbl and $125/bbl. 

The Alliance studied three different cases. The first case applied the 10-ppm sulfur cap 

to RFG. The second case applied the 10-ppm sulfur cap and the 7.0-psi low-RVP standard to 

RFG as well as 7.0- and 7.8-psi low-RVP gasoline. The third case applies the 10-ppm sulfur cap 

and 7.0-psi RVP standard to all RFG and CG. Of these three cases, the first case is most relevant 

because applying the fuels changes to RFG solely applies the 10-ppm sulfur cap to RFG and 

does not involve any changes in RVP. However, the 10-ppm sulfur cap standard studied by the 

Alliance is still 5 ppm more stringent than the 10-ppm average standard that we are proposing. 

The Alliance cost estimate for Case 1 is 1.6 cents per gallon for RFG in PADDs 1, 2 and 

3. This cost estimate is based on amortizing the capital costs on a 10 percent after-tax return on 

investment (ROI). We adjusted the cost estimate to amortize the capital costs based on a before 

tax 7 percent ROI and adjusted the costs to 2010 dollars which increases the costs to 1.75 cents 

per gallon. The 1.75 ¢/gal cost estimate is based on a crude oil price of $51/bbl. The Alliance 

estimated the cost of a 10-ppm sulfur cap standard on RFG assuming that crude oil is priced at 

$125/bbl. At the $125/bbl crude oil price, the Alliance study estimates that it costs 2.50 ¢/gal to 

require that RFG comply with a 10-ppm sulfur cap standard. Adjusting the Alliance costs to 

reflect a 7 percent before tax ROI and 2010 dollars increases the Alliance costs based on a 

$125/bbl crude oil price to 2.69 ¢/gal. 

For our cost analysis we analyzed the cost of sulfur control assuming that crude oil is 

priced at $91.8/bbl. We can interpolate between the Alliance costs based on $51 and $125 per 

barrel crude oil prices, which results in a single cost which is 2.3 cents per gallon. We also 

estimated a cost for refiners lowering their gasoline sulfur to 5 ppm using the refinery-byrefinery cost model and our cost is 1.38 ¢/gal. 

In response to the Alliance study, API retained Baker and O’Brien (BOB) in 2010 to 

study the cost of additional gasoline sulfur control and an increase in RVP using a refinery-byrefinery cost approach with BOB’s Prism model.12

 The Prism model is largely a spreadsheet 

cost model with blending optimization. The primary case analyzed by the API study is the cost 
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of reducing gasoline sulfur to an average of 10 ppm and reducing gasoline RVP to 7.0 psi 

without a 1-psi waiver for blending 10 percent ethanol. The study also analyzed three other 

sensitivity cases: 1) a 5-ppm average gasoline sulfur standard with 7 psi RVP limit on 

conventional gasoline without a 1-psi waiver; 2) a 10-ppm average gasoline sulfur standard and a 

7.8-psi RVP limit on conventional gasoline without a 1-psi waiver; and 3) a 10-ppm average 

gasoline sulfur standard with a 7.8-psi RVP limit on all conventional gasoline with a 1-psi 

waiver. 

In an addendum to its fuels study report released in 2011, API contracted with Baker and 

O’Brien to study a sensitivity case 4, which is a sulfur only case, using its PRISM refinery

model. From our understanding of the study, the study parameters seemed to be about the same 

as the original study, except that API solely studied a 10 ppm average gasoline sulfur standard 

(not including any RVP control), the same sulfur standard that we are proposing. However, API 

also assumed that a 20 ppm cap standard would also be in place which would not allow the 

application of an averaging, banking and trading (ABT) program to optimize refinery

investments and minimize overall costs. 

API made a series of conclusions based on the study. Perhaps the most important 

difference with the original study is that API concluded that not a single refinery would shut 

down as a result of the proposed 10 ppm gasoline sulfur control standard, even though API did 

not study the flexibilities of an ABT program and used excessively high capital costs for a 

grassroots FCC postreater (see below). Like the original study, API did not report average costs, 

but reported the marginal costs for the cost study. Marginal costs reflect the cost of the program 

to the refinery or refineries which would incur the highest costs, assuming that the highest cost 

refineries would set the price (or in this case, the price increase) of gasoline. The report 

concluded that marginal costs after the imposition of a 10 ppm gasoline sulfur program would 

increase the price of gasoline by 6 to 9 cents per gallon in most markets. API did not define how 

its statement “in most markets” would apply to the US gasoline supply. API also did not provide 

any justification why it assumed that the refineries that would experience the highest 

desulfurization cost under Tier 3 would also be the same refineries which sets the gasoline price 

in the gasoline market today. 

Although API did not provide an average gasoline desulfurization cost in its report, we 

could calculate an average cost based on the gasoline volume and total annual costs provided. 

The total cost reported in the report for the 10 ppm average gasoline sulfur standard is 

$2390MM/yr and the non-California gasoline volume is 7343 thousand barrels per day. This 

results in an average per-gallon desulfurization cost of $0.89/bbl or 2.12 c/gal. The difference 

between the average cost and marginal cost (price increase) that API is projecting is profit. API

is projecting that the oil industry would profit from 10 ppm low sulfur standard by the roughly 4 

to 7 cents per gallon difference between the average cost and the two marginal price values. 

That per-gallon profit translates into $4 to $8 billion dollars per year in profit. 

The average cost of the 10 ppm average gasoline sulfur standard was calculated using 

API’s methodology for amortizing capital investments. To facilitate a fairer comparison 

between the API cost study and our cost study, we adjusted the API costs to be on a similar basis 

as our costs. We adjusted the API costs to reflect a before-tax 7 percent return on investment 

(ROI) for capital invested for the hydrotreaters and hydrogen plants instead of the after-tax 15 
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percent ROI used by API. This lowered the API estimated costs from 2.12 c/gal to 1.58 c/gal. 

API’s 1.58 cents per gallon cost is still higher than our 0.89 c/gal cost with an ABT program that 

assumes intercompany trading of credits, and higher than our 0.97 c/gal for the case which 

assumes no ABT program. Thus comparing “apples-to-apples” to the to the extent possible, 

API’s 1.58 c/gal estimated cost for complying with a 10 ppm average gasoline sulfur standard 

compares very favorably with our own cost estimates, and not at all near the 25 c/gal value that 

was sometimes quoted from the first API study. The remaining cost difference between our 

estimated costs and those by API are the capital cost assumptions that API used, as discussed 

below. While little detail is is provided by API about what hardware comprises their 

desulfurization units, the inside battery limits (ISBL) and total capital costs for the FCC 

postreaters and FCC pretreaters are provided in API’s report. API’s FCC pretreaters capital 

costs are consistent with the capital costs that we have used for this unit. However, the FCC 

postreater costs used by API are much higher than what we used and have been used in the past 

by others. API’s capital cost for a grassroots FCC postreater is $228 million for a 35,000 bbl/day

unit, or $6540 per/bbl per day. API’s capital cost includes the outside battery limit (OSBL) 

costs. In contrast, the ISBL capital cost that we used for a grassroots FCC postreater is 

$1500/bbl-day for a 30,000 bbl/day grassroots unit, which increases to $1875/bbl/day when the 

offsite costs are added on. Thus, the API capital costs are about 3 1⁄2 times higher than the capital 

costs that we are using for a grassroots FCC postreater. To check our capital costs, we found 

other capital cost estimates to which we could compare our costs, including the capital costs used 

by the National Petroleum Council when it studied the cost of gasoline desulfurization prior to 

Tier 2. Table 5-35 contains a cost comparison of ISBL, and ISBL and OSBL FCC postreater 

capital costs.

Table 5-35 Capital Cost Comparison

Technology EPA

(Tier 3)

API Mathpro 

(ICCT)

Jacobs Exxon 

Scanfining 

(NPC)

IFP

Prime G 

(NPC)

CDTech 

(NPC)

Phillips 

S-Zorb 

(NPC)

ISBL Capital 

Cost ($/bbl/day)

1500 - - 2440 1045 1410 960 860

ISBL and OSBL 

Capital Cost 

($/bbl/day)

1875 6540 1800 3538 1360 1833 1248 1118

Table 5-35 shows that, compared to the average of the rest of the capital cost estimates, 

the API capital cost for FCC postreater is about four times higher. Compared to the next highest 

cost estimate, which is the FCC postreater capital cost from the Jacobs data base in the Haverly 

refinery cost model that we use,Kthe API capital costs are almost two times higher. 

An important distinction must be made with respect to the severity of desulfurization for 

the capital cost comparison made in Table 5-27. For complying with the Tier 2 gasoline sulfur 

 

K

 The installed capital cost for an FCC postreater from the Jacobs data base was adjusted to current year dollars. 

This estimated installed capital cost is several years old and may not represent Jacobs current cost estimate for a 

FCC postreater.
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5-54

standard (Jacobs and NPC costs), a typical refinery would have installed an FCC postreater to 

desulfurize the FCC naphtha from about 800 ppm down to about 75 ppm, a 725 ppm, or a 91

percent sulfur reduction. In the case of a grassroots postreater that would be installed for Tier 3, 

the postreater would treat FCC naphtha already low in sulfur due to the pretreater installed 

before the FCC unit (these refineries are currently complying with Tier 2 using an FCC 

pretreater). Thus, the new grassroots FCC postreater would only have to reduce the FCC naphtha 

from 100 ppm to 25 ppm, a much smaller 75 ppm or 75 percent sulfur reduction. A grassroots 

FCC postreater installed for Tier 2 would typically remove 10 times more sulfur than one 

installed for Tier 3. This is important because a significant portion of the FCC postreater capital 

cost is devoted to avoiding the recombination reactions which occur when hydrogen sulfide 

concentrations are high and react with the olefins contained in the FCC naphtha. Thus, a 

grassroots FCC postreater installed for Tier 3 would be expected to be significantly lower in 

capital cost compared to a Tier 2 FCC postreater. When API presented the costs to us, they

stated that their grassroots capital costs were based on an actual installation for the Tier 2 

program. This could be one reason why the capital costs used by API for its cost study of the 

Tier 3 program are so high. Another way to assess the API capital cost for the FCC postreaters is 

to compare it to the FCC pretreater cost that API is using. FCC pretreaters are much higher 

pressure units and use more expensive metallurgy than FCC postreaters and, for these two 

reasons, are much more expensive than FCC postreaters on a per-barrel basis. However, API’s 

FCC postreater capital costs are about 50 percent more expensive than its FCC pretreater capital 

costs, which is inconsistent with the design requirements of the units. 

API’s estimated range of capital cost for revamping an FCC postreater is also higher than 

our range of capital cost for revamping an FCC postreater, when assessing the revamped costs as 

a percentage of the capital cost for a grassroots unit. API estimates that revamping an FCC 

postreater would cost 30 to 70 percent of the capital cost for a grassroots FCC unit. Our capital 

cost estimate for revamping an FCC naphtha postreaters from 17 to 50 percent of the capital cost 

for a grassroots FCC postreater, however, most of the revamps are estimated to cost at the lower

end of that range. 

The Emissions Control Technology Association (ECTA) retained personnel within 

Navigant Economics to study the costs of a 10 ppm average gasoline sulfur standard and assess 

the ICCT and API cost studies.13 The authors made a number of conclusions. After reviewing 

both the ICCT and API studies, the authors found that a primary difference in estimated costs 

between the two studies was the capital costs. The authors contacted vendor companies that 

license FCC postreater technologies and surveyed the companies to find out what the capital 

costs are for a FCC postreater. As a result of the survey, the report authors concluded that API’s 

capital costs were too high, and those used in the ICCT study were about right. The authors 

found that Baker and O’Brien has a history of exaggerating the economic impacts of EPA rules, 

citing the costs and other impacts of its analysis of the 2007 on-highway heavy-duty proposed 

rulemaking. The authors concluded that the impact of a 10 ppm gasoline sulfur standard on the 

average refining cost would likely be closer to the 1 cent per gallon estimate by the ICCT study.

Furthermore, the report’s authors also pointed out that the marginal cost analysis conducted by

API did not consider the proposed averaging banking and trading (ABT) program that we were 

expected to propose, which would reduce the marginal costs of the Tier 3 proposed rule. Because 

API’s addendums to its original report came out many months after its original report, we 

originally assessed the most similar case to our proposal from API’s original study. This was 
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sensitivity case 3, which studied a 10-ppm average gasoline sulfur standard with a 7.8-psi RVP 

limit on all conventional gasoline with a 1-psi waiver. One of the most important conclusions by

API with respect to sensitivity case 3 was the projected 623 thousand barrels per day (about 6.5

percent of total gasoline demand) reduction in U.S. gasoline production due to reduced blending 

of light hydrocarbons in response to RVP control and the closure of 4 refineries caused by the 

investment hurdle of complying with both the sulfur and RVP standards evaluated. Because of 

API’s significant projected impacts, which differed from those made by the Alliance and by 

ECTA in their studies, we looked closely at this projected impact. We found that this projection 

for a large shortfall in supply is similar to previous projections made by API for earlier 

rulemakings. Baker and O’brien made a very dire projection for diesel fuel supply when 

analyzing the final highway and proposed nonroad diesel fuel sulfur regulations for the proposed 

nonroad diesel rulemaking back in 2003.14 API’s projection was that 12 refineries would be 

shutdown and that U.S. refiners would exit the diesel market resulting in a shortage of diesel fuel 

supply of 639 thousand barrels per day and to make up the shortfall, the distillate market would 

need to be supplied by imports. Since that projection was made back in 2003 and the highway

diesel fuel program is fully implemented and the nonroad diesel fuel program is mostly

implemented, we can look at the impacts to see how well API’s projection played out. 

First, with respect to imports, the distillate import/export market was fairly stable at about 

200 thousand barrels per day of net imports during the period 2003 to 2005, before the highway

and nonroad programs started. If we add the expected increase in imports estimated by API 

caused by the ULSD highway rule, API projected the imports to increase to 549 thousand barrels 

per day in 2006. When the temporary compliance option of the ULSD highway rule ended in 

2010 and nonroad diesel (excluding locomotive and marine diesel because those requirements 

don’t start until 2012) must comply with the ULSD nonroad diesel rule in 2010, API estimated 

that imports will increase to 829 thousand barrels per day including the baseline imports. 

However, what actually occurred is that imports not only did not increase, but they actually 

decreased to the point that the U.S. became net exporters of 428 thousand barrels per day of 

distillate.15 16 Thus, API was off by over 1200 thousand barrels per day in its estimate of 

distillate production in 2010 by U.S. refiners. Figure 5-9 summarizes the imports and exports of 

diesel fuel and gasoline during the period when the clean fuels regulations were being phased in.
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5-56

Figure 5-9 Petroleum Product Imports and Exports During the Implementation of Clean 

Fuels Regulations

One reason why API was so far off in its projection that U.S. distillate supply would 

decrease significantly is that it had projected that 12 refineries would shut down due to the 

highway and nonroad rulemakings, but, in fact, there were very few refinery shutdowns.17 18

Between 2003 and 2011 when the highway and nonroad diesel fuel ULSD programs were 

phasing in (as well as Tier 2, MSAT2, RFS1 and RFS2), there was a total of 5 refinery

shutdowns of refineries which produce transportation fuels. However, also during this time 

period there were reactivations of three previously shutdown transportation fuel producing 

refineries. Considering both the shutdown refineries and reactivated refineries, there was a net 

shutdown of two transportation fuel producing refinery closures. It is unclear why those 5 

transportation fuel producing refineries shutdown, but there is no evidence to suggest they were 

caused by the diesel rules. There are many factors which cause refinery shutdowns. 

To understand these factors it is necessary to revisit the history of the U.S. refining 

industry.

19

 During the second crude oil embargo which occurred during the Carter 

Administration, subsidies were established that incentivized the construction of many small 

refineries. After crude oil prices dropped, those subsidies ended and many of those refineries 

were not economically competitive with larger refineries. Over time the number of refineries 

decreased and the remaining refineries were expanded to supply the U.S. market. This is part of 

a rationalization that most any industry experiences as the industry matures. Figure 5-10 and 

Table 5-36 summarize the changes in the number of refineries, the average size of refineries, and 

the total production capacity of the U.S. refining industry based on data from the Energy

Information Administration (EIA). 
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Figure 5-10 Refining Industry Statistics

Table 5-36 U.S. Refinery Industry Statistics

Year 1982 - 1992 1993 – 2002 2003 - 2011

Net Change in the Number of 

Operable Refineries

-102 -50 -2

Change in Average Refinery

Size (kbbl/day)

+18,400 +31,800 +11,000

Change in Total U.S. Operating 

Refinery Capacity (MMbbl/day)

-1.2 +1.6 +0.4

Figure 5-10 and Table 5-36 show that there were many refinery closures during the 1980s 

and 1990s, although over time the number of refinery closures was diminishing. Therefore, the 

refinery closures which occurred during the 2003 to 2011 time frame could very well be due to 

this rationalization process reflecting a maturing industry. In fact, examining the trend, the 

number of net refinery shutdowns is virtually flat during the 2003 to 2011 time frame relative to 

the previous periods. This is clearly a rationalization of the refining industry because as 

refineries were closing, the average refinery size increased over each of the periods. Finally, the 

total U.S. refinery capacity decreased and increased back to its 1982 levels during the 1980s and 

early 90s (decrease was likely due to the corporate average fuel economy standards), but showed 

a modest increase from the late 1990’s through 2011. 

5.4 Projected Energy Impacts and Impacts on Permitting

Our refinery-by-refinery model was also used to determine the impact the Tier 3 

standards would have on the energy related CO2 emissions and permitting of existing refineries. 

While the Tier 3 proposal will reduce emissions from vehicles, the addition of grass roots units 
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5-58

and revamping of existing units which we project will happen as a result of the Tier 3 sulfur 

standards are likely to result in some increased emissions of regulated air pollutants at refineries. 

Refinery projects designed to meet the new fuel standards could trigger preconstruction air 

permitting requirements under the Clean Air Act and EPA’s New Source Review (NSR) 

regulations. To address this concern, we used our refinery-by-refinery model to estimate the 

likely process and equipment changes that may be required to meet the Tier 3 gasoline standards. 

This information was submitted to EPA’s Office of Air Quality Planning and Standards 

(OAQPS) to provide the inputs that are necessary for the modeling and analysis of the refinery

emissions and permitting impacts of the Tier 3 fuel standards.

Using our refinery-by-refinery model we generated refinery-specific estimates of the 

increased energy, hydrogen, and gasoline octane demands that we estimate will result from the 

proposed Tier 3 standards. We also estimated the increase in sulfur plant recovery unit (SRU) 

loading/operations for the 111 U.S. refineries that we modeled in our analysis. Energy demand 

includes fuel that is needed to generate refinery process heat, steam and electricity. Hydrogen 

demand is associated with increased hydrotreating of Fluid Catalytic Cracking (FCC) naphtha 

and light straight run (LSR) streams. Increased gasoline octane demand results from refineries 

having to replace octane that is lost due to increased FCC naphtha hydrotreating. Increased SRU 

loading results from the increased fuel desulfurization and associated H2S generation. All of 

these incremental demands will be referred to as “demands” in the following sections. We used 

our refinery-by-refinery model to calculate the increase in these various demands for several 

scenarios where sulfur averaging, banking, and trading (ABT) was not allowed and each refinery

had to meet the 10ppm standard, as well as scenarios that allowed ABT between refineries 

owned by the same parent company to minimize the cost of compliance with the Tier 3 

standards.

5.4.1 Emissions Impacts of Different Production Volumes

In addition to considering scenarios with and without ABT we also considered the 

impacts on emissions and permitting of different gasoline production volumes for each refinery. 

In the first case, called the normal case, we considered the incremental demands for each refinery

assuming no change in gasoline production volume. We also considered a case, called the 

maximum demand case, where each refinery maximized gasoline production based on currently 

existing refinery capacity and equipment.

5.4.1.1 Normal Case

The normal case was estimated using each refinery’s predicted yields of FCC naphtha 

and LSR from our refinery-by-refinery model, along with each refinery’s total gasoline 

production volume from EPA’s RFG database. For each refinery the refinery-by-refinery model 

generated specific Tier 3 demands for hydrogen, steam, fuel gas, electricity and gasoline octane 

based on the desulfurization technology used by each refinery for any FCC postreating and LSR 

hydrotreating. To determine the FCC postreating demands the model considers each refinery’s 

volume of FCC naphtha under normal operations, the FCC naphtha sulfur level at the refinery

prior to postreating, and the process requirements of the FCC postreater technology used by that 

refinery. The demands are calculated by multiplying the FCC naphtha volume by the demands 

from the use of the associated FCC postreating technology. Table 5-26 through Table 5-29 show 
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the FCC postreater technology demand averages as applied to refineries on a national basis for 

the 5 and 10 ppm gasoline sulfur standards. Note that the demands vary significantly with the 

FCC naphtha sulfur level prior to postreating.

Similarly, the normal case demands for any LSR blendstocks that require additional 

hydrotreating as a result of the Tier 3 standards were determined based on each refinery’s yield 

of LSR blendstock under normal operations and the demands for the additional LSR 

hydrotreating. These demands, on a national average basis, are listed in Table 5-31. The normal 

case demands for FCC postreating and LSR hydrotreating were then summed to determine the 

increase in energy, hydrogen, and octane demand. To determine the additional sulfur removed 

from gasoline we first calculated the difference between the current gasoline sulfur level of the 

gasoline produced at each refinery according to their compliance reports to EPA and the 

proposed Tier 3 standard. This difference was multiplied by the refinery’s gasoline production 

volume and divided by the number of days of operation to calculate the additional sulfur removal 

level at each refinery. This sulfur removal information was then used to determine the increase 

in SRU loading on a fractional basis by dividing the additional sulfur removal as a result of the 

Tier 3 standards (in tons of sulfur per day) by the refineries SRU process capacity.

5.4.1.2 Maximum Case

We also considered a second demand case, called the maximum case, in which we 

calculated the demands that result from the Tier 3 standards if each refinery maximizes gasoline 

production based on currently existing refinery capacity. For this case we first determined each 

refiners FCC unit process capacity utilization rate in the normal case. The annual FCC unit 

feedstock charge rate for each refinery as reported in the 2009 EIA data was divided by the FCC 

unit design capacity as reported in the Oil and Gas Journal (OGJ) to calculate the capacity 

utilization rate for the normal case. These normal capacity utilization rates were then scaled up 

to reflect maximum capacity utilization rates, and further adjusted using an overdesign factor.

For refineries projected to meet the proposed Tier 3 standards by revamping existing 

FCC postreating units we assumed that their maximum gasoline production rate was equal to the 

rate produced running the FCC unit at 92% of the refinery’s maximum FCC design capacity. 

There were several refineries that are currently operating their FCC unit greater than a 92%

capacity utilization rate. We assumed that these refineries were already operating at their 

maximum annual capacity utilization rate. For refineries projected to install a new FCC 

postreater we similarly assumed that the new unit would be scaled to process the output of the 

FCC unit operating with a 92% utilization rate. For new FCC postreaters, however, we increased 

the results by 15% as an overdesign factor and adjusted the results accordingly. A similar sizing 

approach was taken for refineries we projected would revamp or add new LSR hydrotreating 

capacity to comply with the proposed Tier 3 standards.

The results represent a “maximum” annual gasoline production case for each refinery

under the Tier 3 standards based on each refiners FCC unit design capacity. These cases 

represent a scenario where each refinery projects emissions based on the maximum achievable 

annual production rate for their existing processing units. These cases reflect each refiner’s 

potential emissions impacts as a result of the proposed Tier 3 standards, rather than the existing 
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Tier 2 standards, when operating at maximum FCC rates as opposed to normal operation more 

indicative of national gasoline demand.

5.4.2 Refinery Demand Sourcing

After determining the increased demands for each refinery as a result of the proposed 

Tier 3 regulations we next developed cases for each refinery demand scenario that represented 

different options for sourcing these demands. Some refiners may choose to produce all of the 

required hydrogen and electricity, as well as make up for all of the lost octane at their refinery. 

Others may choose to purchase some or all of the hydrogen, electricity, and high octane 

blendstocks that they would need to comply with the Tier 3 standards from external suppliers. 

These decisions have a significant impact on the emissions and permitting impacts of the 

proposed Tier 3 regulations. In order to bound all possible scenarios we considered both high 

and low impact cases for each refinery demand scenario. In the high impact scenarios we 

assume that each refinery produces all of the required hydrogen, and electricity needs while 

making up for any gasoline octane loss at their refinery. In the low impact scenarios we assume 

that all the necessary hydrogen, electricity, and high octane blendstocks are purchased from an 

external supplier.

In both cases we assumed that fuel gas demands would increase to meet the increased 

thermal demands at the refinery. In the high impact scenarios the refinery’s fuel gas needs 

would be further increased to produce the needed hydrogen and electricity while replacing the 

lost octane in their gasoline. We consulted literature sources to determine the conversion factors 

from MBTU fuel gas to 1,000 standard cubic feet (scf) of hydrogen20 and 1,000 pounds of 

steam21 that are typical for refineries. We also assumed a standard conversion efficiency from 

fuel gas to electricity for our modeling. These conversion factors are shown in Table 5-37

below.

Table 5-37 Fuel Gas Required to Produce Hydrotreater Utilities

Utility Fuel Gas Required (M BTU)

Hydrogen (1,000 SCF) 239

Steam (1,000 lbs) 1530

Electricity (1 kWh) 5.1

We also had to determine the fuel gas demands that would be required to make up for lost 

octane in our high impacts scenario. For this analysis we assumed that all lost octane in the 

gasoline that results from increased FCC postreating to meet the proposed Tier 3 standards 

would be recovered by running the reformers at the refineries at a higher severity as opposed to 

sourcing it from ethanol or other means. This would increase the octane of the reformate and 

offset the octane losses from the FCC naphtha, resulting in no change in the octane of the overall 

gasoline pool. Running the reformer at a higher severity requires a higher reactor processing 

temperature and an increased volumetric rate for the pumps and compressors. These changes 

further increase the fuel gas and electricity demands of the refinery. We estimated these 

demands using information we obtained from our Jacob’s refinery LP modeling analysis. The 

estimates we used for the fuel gas and electricity demands for increasing gasoline octane are 

shown in Table 5-38 below. These demand increases were only used for the high impact 
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scenarios. For the low impact scenarios we assumed that lost octane was recovered by 

purchasing high octane blendstocks resulting in no fuel gas or electricity demand increases at the 

refinery.

Table 5-38 Energy Required to Increase Gasoline Octane

Utility Amount Required to Increase 

Octanea1 Point (per barrel gasoline)

Fuel Gas (M BTU) 1.87

Electricity (kWh) 0.057

Note:

a (R + M)/2 method

5.4.3 Refinery Demand Impacts

Using our refinery-by-refinery model, along with the technology vendor data for new and 

revamped FCC postreating and LSR hydrotreating data (shown in Table 5-26 through Table 5-29

and Table 5-31) we were able to calculate the increases in refinery demands as a result of our 

proposed Tier 3 regulations for each of the various scenarios outlined in the previous sections. 

This information is summarized in Table 5-39 and Table 5-40 below. The refineries have been 

identified by randomly assigned numbers to protect confidential business information (CBI). 

This information was submitted to OAQPS to serve as the basis for their emissions and 

permitting analysis of the Tier 3 regulations. Based on the data provided by EPA, OAQPS 

determined that the increased demands that result from the Tier 3 regulations would exceed the 

permitting thresholds under the New Source Review (NSR) regulations, including the Prevention 

of Significant Deterioration (PSD) and/or the Nonattainment New Source Review (NA NSR) for 

16 refineries in the low impact scenarios and 22 refineries in the high impact scenarios. An 

additional 7 refineries in the low impacts scenarios, and 13 refineries in the high impact 

scenarios, would exceed the threshold to trigger reviews for greenhouse gas emissions only. A 

technical memorandum describing the OAQPS analysis and results is in the public docket for 

this proposal.22
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Table 5-39 Tier 3 Refinery Energy, Hydrogen and Sulfur Plant Demand Increases (10 ppm, 

No ABT Cases)

Refinery 

Number

Demand Estimates

Fuel Gas Demands Sulfur Plant Production Hydrogen

Low Impact Cases 

(Million BTU/Yr)

High Impact Cases 

(Million BTU/Yr) Sulfur 

Production 

Increase 

(Tons 

Sulfur/Day)

Sulfur Plant 

Capacity 

Increase 

(Percent of 

Existing 

Facility)

Hydrogen 

Demand 

Increase 

(million 

scf/year)

Normal 

Volumes

Maximum 

Volumes

Normal 

Volumes

Maximum 

Volumes

1 117,423 117,423 243,241 243,241 0.28 0.02 426

2 0 0 0 0 0.00 0.00 0

3 2,239,110 2,739,637 2,566,562 3,140,286 0.00 0.00 863

4 173,672 173,672 569,009 569,009 0.43 0.03 1,628

5 46,380 58,279 261,528 328,624 0.13 0.04 655

6 4,433 7,291 14,523 23,888 0.04 1.96 42

7 60,923 69,776 199,605 228,609 0.10 0.03 571

8 23,941 29,792 134,995 167,988 0.17 0.07 338

9 107,543 135,777 222,775 281,262 0.20 0.18 390

10 25,924 29,812 146,179 168,106 0.10 0.08 366

11 13,984 16,250 78,855 91,630 0.09 0.05 198

12 76,663 92,786 251,175 303,997 0.43 0.07 719

13 51,432 51,432 168,509 168,509 0.47 1.18 482

14 0 0 6,444 6,444 0.00 0.00 27

15 16,876 16,876 89,463 89,463 0.07 0.13 219

16 831,027 1,244,900 899,341 1,347,235 0.26 0.22 218

17 31,964 31,964 82,442 82,442 2.19 0.96 206

18 150,568 150,568 311,900 311,900 0.22 0.03 547

19 27,087 44,061 152,736 248,452 0.26 a383

20 43,909 46,106 143,861 151,058 0.37 0.08 412

21 165,539 212,719 343,091 440,875 0.17 0.04 538

22 14,886 16,690 38,395 43,048 0.01 0.02 96

23 83,222 83,222 272,663 272,663 0.41 0.12 780

24 190,870 210,535 206,560 227,841 0.33 0.29 50

25 702,808 702,808 798,012 798,012 0.10 0.09 256

26 20,503 22,001 67,175 72,082 0.90 0.60 192

27 0 0 0 0 0.00 0.00 0

28 3,145,454 3,528,532 3,223,899 3,616,532 0.31 0.07 201

29 0 0 0 0 0.00 0.00 0

30 64,080 66,959 165,279 172,703 0.15 0.01 414

31 0 0 0 0 0.00 0.00 0

32 62,002 62,002 159,918 159,918 0.35 3.46 400

33 1,693,620 1,948,219 1,832,841 2,108,370 0.27 0.03 444

34 16,574 24,879 93,456 140,290 0.15 0.21 234

35 10,469 10,469 12,340 12,340 0.04 0.96 6

36 337,961 411,915 365,743 445,776 0.15 a89

37 5,175 5,175 29,180 29,180 0.13 3.30 73

38 0 0 2,907 2,907 0.00 0.00 12
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Refinery 

Number

Demand Estimates

Fuel Gas Demands Sulfur Plant Production Hydrogen

Low Impact Cases 

(Million BTU/Yr)

High Impact Cases 

(Million BTU/Yr)

Sulfur 

Production 

Sulfur Plant 

Capacity 

Hydrogen 

Demand 

39 21,773 21,773 71,336 71,336 0.58 0.93 204

40 19,171 23,883 91,207 113,621 0.04 0.01 214

41 177,093 203,657 636,542 732,024 0.30 0.03 1,288

42 3,288 4,683 10,773 15,343 0.03 0.97 31

43 41,705 43,146 107,567 111,284 0.52 0.26 269

44 3,609,133 4,265,613 3,905,818 4,616,262 0.66 0.22 946

45 27,946 41,431 105,829 156,897 0.06 a221

46 6,846 8,054 38,603 45,416 0.05 0.55 97

47 1,416,139 1,416,139 1,532,550 1,532,550 0.25 0.03 371

48 21,918 41,206 45,403 85,358 0.02 0.01 80

49 7,024 7,191 18,116 18,547 0.19 2.16 45

50 2,170,692 2,170,692 2,349,131 2,349,131 0.61 0.23 569

51 0 0 0 0 0.00 0.00 0

52 3,945,526 3,984,109 4,269,863 4,311,617 0.61 0.03 1,034

53 102,473 124,044 264,305 319,941 0.46 a662

54 40,155 51,299 108,185 138,209 0.11 0.09 236

55 21,405 26,758 120,695 150,884 0.16 0.13 302

56 26,807 28,515 70,005 74,463 0.04 0.13 144

57 91,964 107,033 518,563 603,534 0.00 0.03 1,299

58 0 0 0 0 0.00 0.00 0

59 52,160 62,683 108,050 129,848 0.10 0.03 189

60 109,050 109,050 357,285 357,285 0.62 0.11 1,022

61 23,773 23,773 36,151 36,151 0.04 0.00 52

62 9,233 9,655 30,251 31,632 0.04 1.96 87

63 136,167 175,790 282,068 364,148 0.08 0.01 494

64 54,151 60,794 165,033 185,278 0.24 0.12 221

65 360,695 402,762 390,346 435,870 0.05 0.19 95

66 8,113 8,589 20,926 22,152 0.04 0.39 52

67 26,289 34,157 67,806 88,100 0.19 0.17 170

68 89,955 103,448 507,237 583,323 0.20 0.02 1,271

69 30,259 33,479 99,139 109,690 0.14 0.14 284

70 9,369 20,299 61,521 133,298 0.20 0.03 103

71 37,188 37,188 121,841 121,841 1.31 0.44 349

72 1,079,231 1,299,822 1,167,948 1,406,672 0.09 0.14 283

73 15,845 17,618 51,915 57,724 0.03 0.00 149

74 38,525 40,291 101,764 106,429 0.09 0.04 215

75 0 0 0 0 0.00 0.00 0

76 14,588 14,588 47,795 47,795 0.03 0.00 137

77 16,875 20,135 43,524 51,932 0.13 0.70 109

78 139,558 139,558 289,093 289,093 0.16 0.02 507

79 0 0 0 0 0.00 0.00 0

80 64,517 76,071 65,729 77,500 0.07 0.10 234

81 2,009,982 2,009,982 2,280,122 2,280,122 0.24 0.05 729

82 58,475 60,441 191,583 198,026 0.26 a548
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-64

Refinery 

Number

Demand Estimates

Fuel Gas Demands Sulfur Plant Production Hydrogen

Low Impact Cases 

(Million BTU/Yr)

High Impact Cases 

(Million BTU/Yr)

Sulfur 

Production 

Sulfur Plant 

Capacity 

Hydrogen 

Demand 

83 2,008,357 2,021,700 2,173,451 2,187,891 0.48 0.09 526

84 0 0 0 0 0.00 0.00 0

85 23,472 27,526 132,354 155,214 0.27 0.14 332

86 20,570 24,731 115,988 139,454 0.00 0.00 291

87 97,698 112,353 363,505 418,031 0.35 0.10 752

88 4,171 5,743 27,391 37,710 0.08 0.12 46

89 147,823 147,823 306,331 306,331 0.13 0.01 495

90 786,104 984,013 850,725 1,064,903 0.30 0.18 206

91 786,104 984,013 850,725 1,064,903 0.41 1.47 206

92 25,608 25,608 53,152 53,152 0.03 0.08 56

93 35,993 47,565 36,669 48,459 0.05 a131

94 4,151,496 4,231,505 4,606,346 4,695,121 0.65 0.08 1,307

95 1,232 2,692 8,090 17,675 0.01 0.03 14

96 2,457 3,308 13,417 18,060 0.02 a33

97 28,695 28,695 74,012 74,012 0.52 a185

98 0 0 0 0 0.00 0.00 0

99 46,468 50,408 152,246 165,155 0.26 0.66 436

100 2,591,010 2,708,025 2,804,001 2,930,635 0.83 0.13 679

101 0 0 0 0 0.00 0.00 0

102 32,339 57,437 111,894 198,734 0.12 0.06 221

103 0 0 0 0 0.00 0.00 0

104 74,538 74,538 154,710 154,710 0.00 0.00 162

105 5,400 5,668 35,458 37,218 0.00 0.00 59

106 139,029 139,029 455,507 455,507 0.25 0.02 1,303

107 0 0 11,540 11,540 0.05 a48

108 8,314 8,879 17,223 18,393 0.00 0.00 30

109 1,485,414 1,704,414 1,607,521 1,844,524 0.32 0.05 389

110 122,762 122,762 279,765 279,765 0.22 0.06 515

111 27,548 33,475 155,338 188,757 0.10 0.07 389

a The refinery did not have published information on the capacity of the existing sulfur plant.
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-65

Table 5-40 Tier 3 Refinery Energy, Hydrogen and Sulfur Plant Demand Increase

(10 ppm, ABT Cases)

Refinery 

Number

Demand Estimates

Fuel Gas Demands Sulfur Plant Production Hydrogen

Low Impact Cases 

(Million BTU/Yr)

High Impact Cases 

(Million BTU/Yr) Sulfur 

Production 

Increase 

(Tons 

Sulfur/Day)

Sulfur Plant 

Capacity 

Increase 

(Percent of 

Existing 

Facility)

Hydrogen 

Demand 

Increase 

(million 

Normal scf/day)

Volumes

Maximum 

Volumes

Normal 

Volumes

Maximum 

Volumes

1 20,503 22,001 67,175 72,082 0.90 0.60 192

2 64,080 66,959 229,725 240,045 0.22 0.02 467

3 323,504 388,955 476,338 572,710 0.00 0.00 528

4 89,955 103,448 507,237 583,323 0.20 0.02 1,271

5 0 0 0 0 0.00 0.00 0

6 426,001 692,966 627,257 1,020,346 0.29 a696

7 0 0 0 0 0.00 0.00 0

8 26,807 28,515 70,005 74,463 0.04 0.13 144

9 117,423 117,423 243,241 243,241 0.28 0.02 426

10 5,400 5,668 35,458 37,218 0.00 0.00 59

11 1,525,387 1,525,387 1,558,831 1,558,831 0.51 1.28 64

12 8,113 8,589 20,926 22,152 0.04 0.39 52

13 7,024 7,191 18,116 18,547 0.19 2.16 45

14 74,538 74,538 154,710 154,710 0.00 0.00 162

15 0 0 0 0 0.00 0.00 0

16 2,591,010 2,708,025 2,804,001 2,930,635 0.83 0.13 679

17 177,093 203,657 636,542 732,024 0.30 0.03 1,288

18 0 0 0 0 0.00 0.00 0

19 19,171 23,883 91,207 113,621 0.04 0.01 214

20 26,289 34,157 67,806 88,100 0.19 0.17 170

21 2,239,110 2,739,637 2,566,562 3,140,286 0.00 0.00 863

22 83,222 83,222 272,663 272,663 0.41 0.12 780

23 147,823 147,823 306,331 306,331 0.13 0.01 495

24 0 0 2,907 2,907 0.00 0.00 12

25 2,273,706 2,751,864 2,323,556 2,812,199 0.50 0.08 96

26 0 0 0 0 0.00 0.00 0

27 107,543 135,777 222,775 281,262 0.20 0.18 390

28 16,875 20,135 43,524 51,932 0.13 0.70 109

29 150,568 150,568 311,900 311,900 0.22 0.03 547

30 0 0 0 0 0.00 0.00 0

31 30,259 33,479 99,139 109,690 0.14 0.14 284

32 0 0 0 0 0.00 0.00 0

33 139,558 139,558 289,093 289,093 0.16 0.02 507

34 9,369 20,299 61,521 133,298 0.20 0.03 103

35 0 0 6,444 6,444 0.00 0.00 27

36 0 0 0 0 0.00 0.00 0

37 109,050 109,050 357,285 357,285 0.62 0.11 1,022

38 0 0 0 0 0.00 0.00 0
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-66

Refinery 

Number

Demand Estimates

Fuel Gas Demands Sulfur Plant Production Hydrogen

Low Impact Cases 

(Million BTU/Yr)

High Impact Cases 

(Million BTU/Yr)

Sulfur 

Production 

Sulfur Plant 

Capacity 

Hydrogen 

Demand 

39 10,469 10,469 12,340 12,340 0.04 0.96 6

40 210,772 269,511 319,729 408,832 0.07 0.00 379

41 173,672 173,672 569,009 569,009 0.43 0.03 1,628

42 21,773 21,773 71,336 71,336 0.58 0.93 204

43 3,145,454 3,528,532 3,223,899 3,616,532 0.31 0.07 201

44 337,961 411,915 365,743 445,776 0.15 a89

45 3,945,526 3,984,109 4,269,863 4,311,617 0.61 0.03 1,034

46 97,698 112,353 363,505 418,031 0.35 0.10 752

47 1,416,139 1,416,139 1,532,550 1,532,550 0.25 0.03 371

48 16,876 16,876 89,463 89,463 0.07 0.13 219

49 433,258 526,468 637,943 775,189 0.14 0.09 707

50 4,433 7,291 14,523 23,888 0.04 1.96 42

51 58,475 60,441 191,583 198,026 0.26 a548

52 2,170,692 2,170,692 2,349,131 2,349,131 0.61 0.23 569

53 8,497 8,497 12,247 12,247 0.00 0.00 16

54 31,964 31,964 82,442 82,442 2.19 0.96 206

55 336,635 420,834 495,673 619,649 0.19 0.15 550

56 32,339 57,437 111,894 198,734 0.12 0.06 221

57 38,525 40,291 101,764 106,429 0.09 0.04 215

58 54,151 60,794 165,033 185,278 0.24 0.12 221

59 0 0 0 0 0.00 0.00 0

60 23,773 23,773 36,151 36,151 0.04 0.00 52

61 190,870 210,535 206,560 227,841 0.33 0.29 50

62 3,288 4,683 10,773 15,343 0.03 0.97 31

63 0 0 0 0 0.00 0.00 0

64 2,008,357 2,021,700 2,173,451 2,187,891 0.48 0.09 526

65 2,457 3,308 13,417 18,060 0.02 a33

66 41,705 43,146 149,509 154,676 0.63 0.31 304

67 46,468 50,408 152,246 165,155 0.26 0.66 436

68 1,079,231 1,299,822 1,167,948 1,406,672 0.09 0.14 283

69 3,071 3,325 7,294 7,896 0.00 0.00 17

70 4,123,377 4,123,377 4,213,781 4,213,781 0.38 0.03 174

71 786,104 984,013 850,725 1,064,903 0.30 0.18 206

72 37,188 37,188 121,841 121,841 1.31 0.44 349

73 0 0 0 0 0.00 0.00 0

74 40,155 51,299 108,185 138,209 0.11 0.09 236

75 1,485,414 1,704,414 1,607,521 1,844,524 0.32 0.05 389

76 122,762 122,762 279,765 279,765 0.22 0.06 515

77 102,473 124,044 264,305 319,941 0.46 a662

78 0 0 0 0 0.00 0.00 0

79 3,609,133 4,265,613 3,905,818 4,616,262 0.66 0.22 946

80 43,909 46,106 143,861 151,058 0.37 0.08 412

81 1,693,620 1,948,219 1,832,841 2,108,370 0.27 0.03 444

82 0 0 0 0 0.00 0.00 0
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-67

Refinery 

Number

Demand Estimates

Fuel Gas Demands Sulfur Plant Production Hydrogen

Low Impact Cases 

(Million BTU/Yr)

High Impact Cases 

(Million BTU/Yr)

Sulfur 

Production 

Sulfur Plant 

Capacity 

Hydrogen 

Demand 

83 60,923 69,776 199,605 228,609 0.10 0.03 571

84 4,171 5,743 27,391 37,710 0.08 0.12 46

85 4,151,496 4,231,505 4,606,346 4,695,121 0.65 0.08 1,307

86 8,314 8,879 17,223 18,393 0.00 0.00 30

87 5,175 5,175 29,180 29,180 0.13 3.30 73

88 96,113 113,326 98,454 116,085 0.13 0.17 543

89 988,379 1,480,617 1,109,185 1,661,587 0.31 0.26 412

90 0 0 0 0 0.00 0.00 0

91 9,233 9,655 30,251 31,632 0.04 1.96 87

92 21,918 41,206 45,403 85,358 0.02 0.01 80

93 14,886 16,690 38,395 43,048 0.01 0.02 96

94 35,993 47,565 36,669 48,459 0.05 a131

95 202,851 261,880 506,929 654,442 0.15 0.01 1,146

96 786,104 984,013 850,725 1,064,903 0.41 1.47 206

97 25,924 29,812 146,179 168,106 0.10 0.08 366

98 1,232 2,692 8,090 17,675 0.01 0.03 14

99 369,152 432,912 543,551 637,434 0.30 0.15 603

100 0 0 0 0 0.00 0.00 0

101 28,695 28,695 74,012 74,012 0.52 a185

102 828,747 828,747 965,965 965,965 0.13 0.11 412

103 0 0 11,540 11,540 0.06 a48

104 13,984 16,250 78,855 91,630 0.09 0.05 198

105 62,002 62,002 159,918 159,918 0.35 3.46 400

106 260,661 391,286 383,805 576,142 0.17 0.25 426

107 194,048 194,048 294,359 294,359 0.07 0.00 349

108 0 0 0 0 0.00 0.00 0

109 428,992 479,023 481,425 537,572 0.08 0.27 179

110 52,160 62,683 108,050 129,848 0.10 0.03 189

111 0 0 0 0 0.00 0.00 0

a The refinery did not have published information on the capacity of the existing sulfur plant.
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-68

Chapter 5 Appendix

LP Refinery Modeling Output Tables
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-69

Table 5-41 Volume and Cost Information Used for Estimating the Cost of Octane

PADD PADD 1 PADD 2 PADD 3 PADD 4/5 USA

Summer 33,001 90,996 248,651 20,296 392,943

Winter 33,692 137,617 302,644 29,719 503,671

Total 66,693 228,613 551,294 50,015 896,614

Reference 10,021,915 19,715,175 51,232,848 11,926,884 92,896,822

Control 9,961,128 19,662,097 51,072,893 11,913,432 92,609,550

Difference -287,272

($/bbl) -0.3203965

(c/gal) -0.7628488

FCC Naphtha 

Volume (kbbl)

Cost 

Obj Function 

($000)
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-70

Table 5-42 PADD 1 Unit Capacity and Throughput Volumes from LP Refinery Modeling 

(Thousand bbl/day)

PADD 1

Refinery Units Summer Winter Summer Winter Summer Winter

Crude Tower 1,260 1,982 2,791 1,982 2,792 0 1

Vacuum Tower 492 788 1,159 788 1,159 0 0

Sats Gas Plant 0 110 120 109 120 -1 0

Unsats Gas Plant 0 157 93 157 93 0 0

FCC DeC5 Tower 0 0 0 0 0 0 0

FCC 490 490 291 490 291 0 0

FCC Splitter 0 0 0 21 0 21 0

Hydrocracker 22 69 69 69 69 0 0

H-Oil Unit 0 0 0 0 0 0 0

Delayed Coker 44 153 292 153 292 0 0

Visbreaker 0 0 0 0 0 0 0

Thermal Naphtha Splitter 0 18 35 18 35 0 0

CRU Reformer 223 281 337 284 337 3 0

SRU Reformer 0 0 0 0 0 0 0

BTX Reformer 0 29 36 25 36 -4 0

C4 Isomerization 32 32 32 32 32 0 0

C5/C6 Isomerization 7 126 149 133 158 7 9

HF Alkylation 40 40 20 40 20 0 0

H2SO4 Alkylation 29 47 44 49 44 3 0

Dimersol 0 0 0 0 0 0 0

Cat Poly 5 1 1 1 1 0 0

Isooctane 0 0 0 0 0 0 0

DHT - Total 334 299 530 299 530 0 0

DHT 2nd RCT - Total 0 240 460 240 460 0 0

DHT Arom Saturation 0 0 0 0 0 0 0

NHT - Total Fd 274 571 650 590 651 19 0

CGH - Generic 119 158 158 158 158 0 0

CGH - Olefin Sat'n 119 0 0 0 0 0 0

FCCU Fd HDT 44 1,459 1,459 1,459 1,459 0 0

LSR Splitter 0 87 51 100 50 13 -1

LSR Bz Saturator 0 0 0 0 0 0 0

Reformate Saturator 0 1 1 1 1 0 0

Reformate Splitter 0 3 3 3 3 0 0

SDA 32,600 33 33 33 33 0 0

MTBE 3 0 0 0 0 0 0

TAME 0 0 0 0 0 0 0

Hydrogen Plant - Total MSCF 64 974 1,007 974 1,008 0 1

Lube Unit 146 53 53 53 53 0 0

Sulfur Plant 1,434 1 2 1 2 0 0

Merox Jet 0 0 0 0 0 0 0

Merox Diesel 0 0 0 0 0 0 0

BTX Reformer - Tower feed 13 11 13 11 13 0 0

BTX Reformer - Extract feed 23 23 23 23 23 0 0

Toluene Dealkyation 0 0 0 0 0 0 0

Cumene 0 0 0 0 0 0 0

Cyclohexane 0 0 0 0 0 0 0

Control Case Minus 1

ON

Control Case Relative

to Ref Case

Actual 2017 Ref case

Capacity 

in 2009
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-71

Table 5-43 PADD 2 Unit Capacity and Throughput Volumes from LP Refinery Modeling 

(Thousand bbl/day)

PADD 2

Refinery Units Summer Winter Summer Winter Summer Winter

Crude Tower 3,612 2,996 3,831 2,996 3,840 0 10

Vacuum Tower 1,452 1,326 1,660 1,326 1,664 0 4

Sats Gas Plant 0 167 171 169 175 2 4

Unsats Gas Plant 0 255 301 255 302 0 1

FCC DeC5 Tower 0 0 0 0 0 0 0

FCC 1,144 852 1,036 852 1,039 0 3

FCC Splitter 0 500 594 500 596 0 2

Hydrocracker 234 216 216 216 216 0 0

H-Oil Unit 0 0 0 0 0 0 0

Delayed Coker 375 280 329 280 329 0 0

Visbreaker 0 0 0 0 0 0 0

Thermal Naphtha Splitter 0 34 40 34 40 0 0

CRU Reformer 807 516 529 540 542 24 13

SRU Reformer 0 0 0 0 0 0 0

BTX Reformer 0 42 54 42 54 0 0

C4 Isomerization 15 65 60 65 63 0 3

C5/C6 Isomerization 157 168 125 170 141 2 16

HF Alkylation 156 126 78 126 78 0 0

H2SO4 Alkylation 81 81 40 81 40 0 0

Dimersol 3 0 0 0 0 0 0

Cat Poly 4 0 0 0 0 0 0

Isooctane 0 0 0 0 0 0 0

DHT - Total 1,056 691 577 691 577 0 0

DHT 2nd RCT - Total 0 512 528 512 529 0 0

DHT Arom Saturation 6 0 0 0 0 0 0

NHT - Total Fd 1,038 818 929 823 932 5 3

CGH - Generic 282 127 176 133 177 5 1

CGH - Olefin Sat'n 282 0 0 0 0 0 0

FCCU Fd HDT 521 521 521 521 521 0 0

LSR Splitter 0 71 100 75 100 4 0

LSR Bz Saturator 0 0 0 0 0 0 0

Reformate Saturator 0 1 1 1 1 0 0

Reformate Splitter 0 3 3 3 3 0 0

SDA 32,600 0 36 0 36 0 0

MTBE 0 0 0 0 0 0 0

TAME 0 0 0 0 0 0 0

Hydrogen Plant - Total MSCF 799,200 1,004 1,066 985 1,058 -19 -7

Lube Unit 66 24 24 24 24 0 0

Sulfur Plant 1,434 4 5 4 5 0 0

Merox Jet 0 0 0 0 0 0 0

Merox Diesel 0 0 0 0 0 0 0

BTX Reformer - Tower feed 18 16 18 16 18 0 0

BTX Reformer - Extract feed 34 34 34 34 34 0 0

Toluene Dealkyation 0 0 0 0 0 0 0

Cumene 0 0 0 0 0 0 0

Cyclohexane 0 0 0 0 0 0 0

2017 Ref case Control Case Minus 1

ON

Control Case Relative

to Ref Case

Actual 

Capacity 

in 2009
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-72

Table 5-44 PADD 3 Unit Capacity and Throughput Volumes from LP Refinery Modeling 

(Thousand bbl/day)

PADD 3

Refinery Units Summer Winter Summer Winter Summer Winter

Crude Tower 8,735 5,576 5,706 5,589 5,706 13 0

Vacuum Tower 3,963 2,587 2,647 2,593 2,647 6 0

Sats Gas Plant 0 534 522 541 523 8 0

Unsats Gas Plant 0 839 728 839 730 0 3

FCC DeC5 Tower 0 0 0 0 0 0 0

FCC 2,808 2,808 2,411 2,808 2,422 0 11

FCC Splitter 0 1,636 1,207 1,636 1,218 0 11

Hydrocracker 523 1,416 1,416 1,397 1,397 -19 -19

H-Oil Unit 0 144 144 144 144 0 0

Delayed Coker 1,350 625 648 628 649 2 2

Visbreaker 51 0 51 0 49 0 -2

Thermal Naphtha Splitter 0 74 79 74 79 0 0

CRU Reformer 1,765 524 410 613 454 89 44

SRU Reformer 0 0 0 0 0 0 0

BTX Reformer 0 381 383 385 383 4 0

C4 Isomerization 50 133 133 133 133 0 0

C5/C6 Isomerization 199 47 0 75 0 28 0

HF Alkylation 295 235 148 235 148 0 0

H2SO4 Alkylation 290 290 267 290 270 0 3

Dimersol 19 0 0 0 0 0 0

Cat Poly 23 0 0 0 0 0 0

Isooctane 0 0 0 0 0 0 0

DHT - Total 2,606 2,080 2,032 2,086 2,038 5 6

DHT 2nd RCT - Total 0 1,592 1,565 1,599 1,570 6 6

DHT Arom Saturation 4 0 0 0 0 0 0

NHT - Total Fd 2,284 1,177 1,041 1,206 1,049 30 8

CGH - Generic 905 173 18 196 26 23 8

CGH - Olefin Sat'n 905 0 0 0 0 0 0

FCCU Fd HDT 1,375 1,559 1,559 1,559 1,559 0 0

LSR Splitter 0 1 1 1 1 0 0

LSR Bz Saturator 0 0 0 0 0 0 0

Reformate Saturator 0 10 1 10 1 0 0

Reformate Splitter 0 30 3 30 3 0 0

SDA 32,600 263 263 263 263 0 0

MTBE 28 0 0 0 0 0 0

TAME 3 0 0 0 0 0 0

Hydrogen Plant - Total MSCF 1,569,400 5,345 5,487 5,217 5,414 -129 -73

Lube Unit 1,169 506 506 506 506 0 0

Sulfur Plant 1,434 15 13 15 13 0 0

Merox Jet 0 0 0 0 0 0 0

Merox Diesel 0 0 0 0 0 0 0

BTX Reformer - Tower feed 146 146 146 146 146 0 0

BTX Reformer - Extract feed 318 318 318 318 318 0 0

Toluene Dealkyation 0 0 0 0 0 0 0

Cumene 0 0 0 0 0 0 0

Cyclohexane 0 0 0 0 0 0 0

Actual

Capacity 

in 2009

2017 Ref case Control Case Minus 1

ON

Control Case Relative

to Ref Case
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-73

Table 5-45 PADD 4 and PADD 5OC Unit Capacity and Throughput Volumes from LP 

Refinery Modeling (Thousand bbl/day)

PADD 4 and 5OC

Refinery Units Summer Winter Summer Winter Summer Winter

Crude Tower 1,796 1,518 1,558 1,518 1,558 0 0

Vacuum Tower 599 559 541 559 541 0 0

Sats Gas Plant 0 71 71 71 68 0 -3

Unsats Gas Plant 0 66 70 66 70 0 0

FCC DeC5 Tower 0 0 0 0 0 0 0

FCC 344 224 224 224 224 0 0

FCC Splitter 0 132 140 132 140 0 0

Hydrocracker 91 144 194 144 194 0 0

H-Oil Unit 0 0 0 0 0 0 0

Delayed Coker 143 133 133 133 133 0 0

Visbreaker 56 41 20 41 20 0 0

Thermal Naphtha Splitter 0 17 16 17 16 0 0

CRU Reformer 0 0 0 0 0 0 0

SRU Reformer 287 214 211 236 248 0 37

BTX Reformer 0 0 0 0 0 0 0

C4 Isomerization 19 19 19 19 19 0 0

C5/C6 Isomerization 43 2 0 2 0 0 0

HF Alkylation 37 19 36 19 36 0 0

H2SO4 Alkylation 38 38 38 38 38 0 0

Dimersol 1 1 0 1 0 0 0

Cat Poly 10 0 0 0 0 0 0

Isooctane 0 0 0 0 0 0 0

DHT - Total 404 492 483 492 483 0 0

DHT 2nd RCT - Total 0 331 331 331 331 0 0

DHT Arom Saturation 0 0 0 0 0 0 0

NHT - Total Fd 326 426 435 434 434 0 -1

CGH - Generic 66 0 0 0 0 0 0

CGH - Olefin Sat'n 66 0 0 0 0 0 0

FCCU Fd HDT 89 89 89 89 89 0 0

LSR Splitter 0 100 100 100 100 0 0

LSR Bz Saturator 0 0 0 0 0 0 0

Reformate Saturator 0 10 4 1 1 0 -3

Reformate Splitter 0 30 11 3 3 0 -8

SDA 32,600 14 28 14 28 0 0

MTBE 0 0 0 0 0 0 0

TAME 0 0 0 0 0 0 0

Hydrogen Plant - Total MSCF 373,900 518 619 499 603 0 -16

Lube Unit 0 0 0 0 0 0 0

Sulfur Plant 1,434 2 2 2 2 0 0

Merox Jet 0 0 0 0 0 0 0

Merox Diesel 0 0 0 0 0 0 0

BTX Reformer - Tower feed 0 0 0 0 0 0 0

BTX Reformer - Extract feed 0 0 0 0 0 0 0

Toluene Dealkyation 0 0 0 0 0 0 0

Cumene 0 0 0 0 0 0 0

Cyclohexane 0 0 0 0 0 0 0

2017 Ref case Control Case Minus 1

ON

Control Case Relative

to Ref Case

Actual

Capacity 

in 2009
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-74

Table 5-46 U.S. (except CA) Unit Capacity and Throughput Volumes from LP Refinery 

Modeling (Thousand bbl/day)

PADDs 1 -4 and 5OC

Refinery Units Summer Winter Summer Winter Summer Winter

Crude Tower 15,403 12,071 13,886 12,084 13,897 13 11

Vacuum Tower 6,505 5,260 6,007 5,266 6,012 6 5

Sats Gas Plant 0 882 885 890 886 9 1

Unsats Gas Plant 0 1,317 1,192 1,317 1,196 0 3

FCC DeC5 Tower 0 0 0 0 0 0 0

FCC 4,786 4,373 3,962 4,373 3,976 0 13

FCC Splitter 0 2,267 1,941 2,288 1,954 20 12

Hydrocracker 869 1,845 1,895 1,826 1,876 -20 -20

H-Oil Unit 0 144 144 144 144 0 0

Delayed Coker 1,912 1,190 1,401 1,193 1,403 2 2

Visbreaker 107 41 71 41 70 0 -2

Thermal Naphtha Splitter 0 143 170 144 171 0 0

CRU Reformer 2,794 1,321 1,276 1,437 1,333 115 57

SRU Reformer 287 214 211 236 248 0 37

BTX Reformer 0 452 473 452 473 1 0

C4 Isomerization 116 250 245 249 247 -1 2

C5/C6 Isomerization 406 343 275 380 299 37 25

HF Alkylation 528 420 281 420 281 0 0

H2SO4 Alkylation 438 455 389 458 392 3 3

Dimersol 23 1 0 1 0 0 0

Cat Poly 42 1 1 1 1 0 0

Isooctane 0 0 0 0 0 0 0

DHT - Total 4,400 3,562 3,622 3,567 3,628 5 6

DHT 2nd RCT - Total 0 2,676 2,885 2,682 2,891 6 6

DHT Arom Saturation 10 0 0 0 0 0 0

NHT - Total Fd 3,921 2,992 3,056 3,054 3,066 53 10

CGH - Generic 1,373 459 352 487 361 28 9

CGH - Olefin Sat'n 1,373 0 0 0 0 0 0

FCCU Fd HDT 2,029 3,628 3,628 3,628 3,628 0 0

LSR Splitter 0 259 252 276 251 17 -1

LSR Bz Saturator 0 0 0 0 0 0 0

Reformate Saturator 0 22 7 13 4 0 -3

Reformate Splitter 0 66 20 39 12 0 -8

SDA 130,400 310 359 310 359 0 0

MTBE 31 0 0 0 0 0 0

TAME 3 0 0 0 0 0 0

Hydrogen Plant - Total MSCF 2,742,564 7,841 8,178 7,675 8,084 -147 -95

Lube Unit 1,380 582 582 582 582 0 0

Sulfur Plant 5,736 22 22 22 22 0 0

Merox Jet 0 0 0 0 0 0 0

Merox Diesel 0 0 0 0 0 0 0

BTX Reformer - Tower feed 177 173 177 173 177 0 0

BTX Reformer - Extract feed 375 375 375 375 375 0 0

Toluene Dealkyation 0 0 0 0 0 0 0

Cumene 0 0 0 0 0 0 0

Cyclohexane 0 0 0 0 0 0 0

2017 Ref case Control Case Minus 1

ON

Control Case Relative

to Ref Case

Actual

Capacity

in 2009
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-75

Table 5-47 PADD 1 Gasoline Qualities Estimated by LP Refinery Modeling

CG RFG Pool CG RFG Pool CG RFG Pool CG RFG Pool

Energy (MMBTU/bbl) 4.84 4.96 4.85 4.87 4.82 4.85 4.84 4.95 4.88 4.87 4.81 4.85

Density (lb/bbl) 254.7 261.9 255.3 255.7 258.1 256.7 255.0 261.7 257.6 255.7 258.2 256.7

Sulfur (ppm) 23.6 22.0 23.5 22.2 24.9 23.3 23.4 21.2 22.6 22.6 24.9 23.6

% at 200 45.9 45.5 45.9 48.7 52.2 50.2 45.3 46.2 45.7 48.4 52.0 49.9

% at 300 81.8 75.8 81.2 81.5 79.1 80.5 81.0 76.5 79.2 80.8 78.6 79.9

RVP (psi) 9.11 7.00 8.92 12.15 11.93 12.06 9.11 7.00 8.28 12.17 11.94 12.07

T10 (F) 124 137 125 104 106 105 124 137 129 104 106 105

T50 (F) 208 209 208 202 195 199 209 208 209 203 196 200

T90 (F) 335 362 338 336 347 341 339 359 347 339 349 344

Driveability 1145 1195 1150 1100 1092 1097 1152 1187 1166 1105 1095 1101

Vapor Lock 137.4 148.8 138.4 120.1 120.1 120.1 137.6 148.6 141.9 120.1 120.1 120.1

Aromatics (vol%) 20.8 22.0 20.9 21.5 21.2 21.3 22.1 22.0 22.0 22.3 21.7 22.0

Benzene (vol%) 0.42 0.42 0.42 0.50 0.63 0.55 0.45 0.51 0.47 0.51 0.63 0.56

Olefins (vol%) 9.17 11.48 9.38 8.61 8.30 8.48 9.09 11.35 9.98 8.74 8.15 8.50

Alcohol (vol%) 11.2 13.8 11.4 11.1 13.7 12.2 11.2 13.8 12.2 11.1 13.7 12.2

Oxygen (wt%) 4.23 5.06 4.31 4.19 5.12 4.57 4.22 5.07 4.56 4.19 5.12 4.57

Volumes (kbbl/day) 1,693,907 169,974 1,863,881 345,786 240,886 586,672 1,709,571 1,118,250 2,827,821 1,623,409 1,130,922 2,754,331

Gasoline Qualities

and Volume

2017 Ref Case 2017 minus 1 ON in FCC naphtha

Summer Winter Summer Winter

Table 5-48 PADD 2 Gasoline Qualities Estimated by LP Refinery Modeling

CG RFG Pool CG RFG Pool CG RFG Pool CG RFG Pool

Energy (MMBTU/bbl) 4.91 4.97 4.92 4.89 4.84 4.88 4.91 4.97 4.92 4.88 4.83 4.88

Density (lb/bbl) 259.9 262.4 260.3 259.1 258.4 259.1 260.0 262.1 260.3 259.3 258.4 259.2

Sulfur (ppm) 27.4 20.3 26.3 25.2 22.7 25.1 26.5 18.9 25.3 25.2 22.4 25.0

% at 200 56.2 52.4 55.6 58.0 59.2 58.1 56.0 52.4 55.5 57.9 59.3 58.0

% at 300 82.9 80.0 82.4 81.7 82.9 81.8 82.8 80.2 82.4 81.6 82.9 81.7

RVP (psi) 9.03 7.00 8.72 11.55 11.47 11.55 9.03 7.00 8.72 11.56 11.47 11.55

T10 (F) 124 137 126 108 109 108 124 137 126 108 109 108

T50 (F) 187 195 188 184 181 183 188 195 189 184 181 183

T90 (F) 330 343 332 335 330 335 331 342 332 336 330 335

Driveability 1077 1133 1086 1048 1036 1047 1079 1132 1088 1049 1036 1048

Vapor Lock 134.2 146.4 136.1 120.1 120.1 120.1 134.3 146.4 136.2 120.1 120.1 120.1

Aromatics (vol%) 21.2 21.2 21.2 22.1 20.1 21.9 21.8 21.2 21.7 22.4 20.0 22.2

Benzene (vol%) 0.77 0.88 0.79 0.49 0.67 0.50 0.76 0.88 0.78 0.50 0.65 0.51

Olefins (vol%) 8.66 5.44 8.16 9.10 8.50 9.06 8.83 5.08 8.25 9.14 8.40 9.08

Alcohol (vol%) 11.2 13.8 11.6 11.2 12.0 11.3 11.2 13.8 11.6 11.2 12.0 11.3

Oxygen (wt%) 4.17 5.06 4.31 4.18 4.48 4.20 4.16 5.07 4.31 4.18 4.48 4.20

Volumes (kbbl/day) 1,799,149 334,626 2,133,775 1,710,734 139,111 1,849,845 1,799,149 334,626 2,133,775 1,710,734 139,111 1,849,845

Gasoline Qualities

and Volume

2017 Ref Case 2017 minus 1 ON in FCC naphtha

Summer Winter Summer Winter
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-76

Table 5-49 PADD 3 Gasoline Qualities Estimated by LP Refinery Modeling

CG RFG Pool CG RFG Pool CG RFG Pool CG RFG Pool

Energy (MMBTU/bbl) 5.06 4.90 5.01 4.99 4.98 4.98 5.02 5.00 5.01 4.99 4.96 4.98

Density (lb/bbl) 259.5 254.7 258.2 254.8 255.8 255.1 258.6 261.2 259.3 255.0 255.6 255.2

Sulfur (ppm) 28.0 27.2 27.8 28.0 28.0 28.0 28.0 27.2 27.8 28.0 28.0 28.0

% at 200 59.0 59.4 59.1 59.6 61.4 60.1 59.5 51.8 57.4 59.3 62.0 60.0

% at 300 83.6 88.2 84.9 85.1 85.3 85.2 84.0 80.0 82.9 84.6 85.4 84.8

RVP (psi) 9.07 7.00 8.49 11.45 11.33 11.42 9.07 7.00 8.49 11.47 11.29 11.42

T10 (F) 124 137 128 109 110 109 124 137 128 109 110 109

T50 (F) 182 181 181 180 177 179 180 196 185 181 175 179

T90 (F) 327 306 321 320 319 320 325 343 330 322 319 321

Driveability 1057 1053 1056 1024 1013 1021 1052 1137 1075 1028 1009 1023

Vapor Lock 133.1 144.0 136.1 120.1 120.1 120.1 132.9 146.6 136.7 120.1 120.1 120.1

Aromatics (vol%) 20.8 6.5 16.9 15.0 15.2 15.1 19.8 18.3 19.4 15.6 15.3 15.5

Benzene (vol%) 0.62 0.61 0.61 0.58 0.52 0.56 0.61 0.61 0.61 0.58 0.61 0.58

Olefins (vol%) 13.42 1.72 10.18 10.82 10.90 10.84 13.07 3.17 10.33 10.90 10.82 10.88

Alcohol (vol%) 11.2 13.7 11.9 11.2 13.7 11.8 11.2 13.7 11.9 11.2 13.7 11.8

Oxygen (wt%) 4.15 5.20 4.44 4.23 5.17 4.47 4.16 5.07 4.42 4.22 5.18 4.47

Volumes (kbbl/day) 1,276,784 488,041 1,764,825 1,400,642 497,125 1,897,767 1,276,784 488,041 1,764,825 1,400,642 497,125 1,897,767

Gasoline Qualities

and Volume

2017 Ref Case 2017 minus 1 ON in FCC naphtha

Summer Winter Summer Winter

Table 5-50 PADD 4 and 5OC Gasoline Qualities Estimated by LP Refinery Modeling

CG RFG Pool CG RFG Pool CG RFG Pool CG RFG Pool

Energy (MMBTU/bbl) 4.96 0.00 4.96 4.95 0.00 4.95 4.96 0.00 4.96 4.95 0.00 4.95

Density (lb/bbl) 263.3 0.0 263.3 261.2 0.0 261.2 263.5 0.0 263.5 261.1 0.0 261.1

Sulfur (ppm) 25.0 0.0 25.0 28.0 0.0 28.0 24.6 0.0 24.6 28.0 0.0 28.0

% at 200 58.5 0.0 58.5 58.6 0.0 58.6 59.0 0.0 59.0 57.9 0.0 57.9

% at 300 85.5 0.0 85.5 84.7 0.0 84.7 85.7 0.0 85.7 83.0 0.0 83.0

RVP (psi) 9.34 0.00 9.34 11.52 0.00 11.52 9.34 0.00 9.34 11.55 0.00 11.55

T10 (F) 122 0 122 108 0 108 122 0 122 108 0 108

T50 (F) 182 0 182 182 0 182 181 0 181 184 0 184

T90 (F) 318 0 318 322 0 322 317 0 317 329 0 329

Driveability 1048 0 1048 1031 0 1031 1045 0 1045 1042 0 1042

Vapor Lock 131.8 0.0 131.8 120.0 0.0 120.0 131.6 0.0 131.6 120.1 0.0 120.1

Aromatics (vol%) 24.4 0.0 24.4 23.8 0.0 23.8 24.5 0.0 24.5 23.4 0.0 23.4

Benzene (vol%) 0.84 0.00 0.84 1.09 0.00 1.09 0.86 0.00 0.86 1.10 0.00 1.10

Olefins (vol%) 5.70 0.00 5.70 5.47 0.00 5.47 5.72 0.00 5.72 4.98 0.00 4.98

Alcohol (vol%) 11.3 0.0 11.3 11.1 0.0 11.1 11.3 0.0 11.3 11.3 0.0 11.3

Oxygen (wt%) 4.15 0.00 4.15 4.08 0.00 4.08 4.14 0.00 4.14 4.18 0.00 4.18

Volumes (kbbl/day) 701,850 0 701,850 734,023 0 734,023 701,850 0 701,850 744,898 0 744,898

Gasoline Qualities

and Volume

2017 Ref Case 2017 minus 1 ON in FCC naphtha

Summer Winter Summer Winter
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-77

Table 5-51 U.S. (except CA) Gasoline Qualities Estimated by Refinery Modeling

CG RFG Pool CG RFG Pool CG RFG Pool CG RFG Pool

Energy (MMBTU/bbl) 4.93 4.93 4.93 4.93 4.91 4.93 4.92 4.96 4.93 4.91 4.85 4.90

Density (lb/bbl) 258.6 258.5 258.6 257.8 256.9 257.6 258.6 261.7 259.4 257.4 257.5 257.4

Sulfur (ppm) 26.1 24.0 25.8 26.4 26.3 26.4 25.7 22.3 24.8 25.5 25.6 25.5

% at 200 54.0 54.7 54.1 57.9 58.5 58.0 53.9 48.7 52.5 55.4 55.4 55.4

% at 300 83.1 83.3 83.1 83.3 83.2 83.3 82.9 78.0 81.6 82.3 80.8 82.0

RVP (psi) 9.10 7.00 8.78 11.56 11.52 11.55 9.10 7.00 8.55 11.71 11.72 11.72

T10 (F) 124 137 126 108 108 108 124 137 127 107 107 107

T50 (F) 192 190 191 184 182 184 192 202 195 189 189 189

T90 (F) 329 328 329 328 328 328 330 352 336 333 339 334

Driveability 1090 1104 1092 1041 1038 1041 1091 1165 1111 1059 1066 1061

Vapor Lock 134.6 145.7 136.3 120.1 120.1 120.1 134.6 147.7 138.1 120.1 120.1 120.1

Aromatics (vol%) 21.4 14.1 20.3 20.0 17.6 19.6 21.8 20.9 21.5 20.8 19.8 20.5

Benzene (vol%) 0.63 0.67 0.64 0.63 0.57 0.62 0.64 0.60 0.63 0.60 0.62 0.61

Olefins (vol%) 9.55 4.65 8.80 9.00 9.81 9.14 9.50 8.21 9.16 8.91 8.92 8.91

Alcohol (vol%) 11.2 13.8 11.6 11.2 13.4 11.6 11.2 13.8 11.9 11.2 13.6 11.8

Oxygen (wt%) 4.18 5.13 4.33 4.18 5.05 4.33 4.18 5.07 4.41 4.19 5.08 4.41

Volumes (kbbl/day) 5,471,690 992,641 6,464,331 4,191,185 877,122 5,068,307 5,487,354 1,940,917 7,428,271 5,479,683 1,767,158 7,246,841

Gasoline Qualities

and Volume

2017 Ref Case 2017 minus 1 ON in FCC naphtha

Summer Winter Summer Winter
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-78

Table 5-52 PADD 1 Unit Capacity and Throughput Volumes from LP Refinery Modeling 

(Thousand bbl/day)

PADD 1

Refinery Units Summer Winter Summer Winter Summer Winter Summer Winter Summer Winter

Crude Tower 1,260 1,156 1,716 757 1,267 -400 -449 757 1,267 -400 -449

Vacuum Tower 492 507 738 332 550 -175 -187 332 550 -175 -187

Sats Gas Plant 0 104 62 70 50 -34 -12 72 47 -32 -15

Unsats Gas Plant 0 185 226 156 191 -29 -36 156 191 -29 -36

FCC DeC5 Tower 0 0 0 0 0 0 0 0 0 0 0

FCC 490 678 827 565 691 -113 -136 565 691 -113 -136

FCC Splitter 0 387 467 324 391 -63 -76 324 391 -63 -76

Hydrocracker 22 52 82 52 82 0 0 52 82 0 0

H-Oil Unit 0 0 0 0 0 0 0 0 0 0 0

Delayed Coker 44 22 94 22 82 0 -12 22 82 0 -12

Visbreaker 0 0 0 0 0 0 0 0 0 0 0

Thermal Naphtha Splitter 0 3 11 3 10 0 -1 3 10 0 -1

CRU Reformer 223 138 223 90 154 -48 -70 134 154 -3 -70

SRU Reformer 0 0 0 0 0 0 0 0 0 0 0

BTX Reformer 0 36 36 36 36 0 0 36 36 0 0

C4 Isomerization 32 53 0 32 0 -21 0 32 0 -21 0

C5/C6 Isomerization 7 51 95 32 32 -19 -63 17 8 -34 -87

HF Alkylation 40 40 20 40 20 0 0 40 20 0 0

H2SO4 Alkylation 29 125 68 94 20 -31 -48 77 20 -48 -48

Dimersol 0 0 0 0 0 0 0 0 0 0 0

Cat Poly 5 1 31 1 32 0 1 8 32 7 1

Isooctane 0 0 0 0 0 0 0 0 0 0 0

DHT - Total 334 431 742 300 572 -130 -170 310 572 -121 -170

DHT 2nd RCT - Total 0 0 0 0 0 0 0 0 0 0 0

DHT Arom Saturation 0 0 0 0 0 0 0 0 0 0 0

NHT - Total Fd 274 263 397 165 286 -98 -111 152 286 -111 -111

CGH - Generic 119 161 158 159 158 -2 0 158 158 -3 0

CGH - Olefin Sat'n 119 0 0 0 0 0 0 0 0 0 0

FCCU Fd HDT 44 460 460 396 396 -63 -63 396 396 -63 -63

LSR Splitter 0 1 51 1 24 0 -27 1 40 0 -11

LSR Bz Saturator 0 0 0 0 0 0 0 0 0 0 0

Reformate Saturator 0 21 1 19 1 -2 0 10 1 -11 0

Reformate Splitter 0 63 3 57 3 -6 0 30 3 -33 0

SDA 0 33 33 33 33 0 0 33 33 0 0

MTBE 3 0 0 0 0 0 0 0 0 0 0

TAME 0 0 0 0 0 0 0 0 0 0 0

Hydrogen Plant - Total MSCF 63,700 645 935 538 788 -107 -147 518 795 -127 -140

Lube Unit 146 53 53 53 53 0 0 53 53 0 0

Sulfur Plant 1,434 1 1 1 1 0 0 1 1 0 0

Merox Jet 0 0 0 0 0 0 0 0 0 0 0

Merox Diesel 0 0 0 0 0 0 0 0 0 0 0

BTX Reformer - Tower feed 13 13 13 13 13 0 0 13 13 0 0

BTX Reformer - Extract feed 23 23 23 23 23 0 0 23 23 0 0

Toluene Dealkyation 0 0 0 0 0 0 0 0 0 0 0

Cumene 0 0 0 0 0 0 0 0 0 0 0

Cyclohexane 0 0 0 0 0 0 0 0 0 0 0

Reference 2005 in

2030 

100% E10 Case in

2030

E10 Case Relative to 

Ref Case

100% E15 Case in

2030

E15 Case Relative to

Ref Case

Actual 

Capacity

in 2009
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-79

Table 5-53 PADD 2 Unit Capacity and Throughput Volumes from LP Refinery Modeling 

(Thousand bbl/day)

PADD 2

Refinery Units Summer Winter Summer Winter Summer Winter Summer Winter Summer Winter

Crude Tower 3,612 2,638 4,141 2,494 3,972 -145 -169 2,206 3,830 -432 -311

Vacuum Tower 1,452 1,092 1,770 1,026 1,711 -66 -58 905 1,656 -186 -114

Sats Gas Plant 0 162 137 136 144 -26 7 107 120 -55 -18

Unsats Gas Plant 0 248 332 204 319 -44 -13 166 302 -82 -31

FCC DeC5 Tower 0 0 0 0 0 0 0 0 0 0 0

FCC 1,144 864 1,144 666 1,080 -197 -64 569 1,031 -295 -113

FCC Splitter 0 503 653 6 628 -496 -24 0 595 -503 -58

Hydrocracker 234 234 234 234 234 0 0 234 234 0 0

H-Oil Unit 0 0 0 0 0 0 0 0 0 0 0

Delayed Coker 375 153 395 129 375 -24 -21 86 355 -67 -40

Visbreaker 0 0 0 0 0 0 0 0 0 0 0

Thermal Naphtha Splitter 0 19 48 16 45 -3 -3 10 43 -8 -5

CRU Reformer 807 393 613 376 515 -18 -98 329 496 -64 -116

SRU Reformer 0 0 0 0 0 0 0 0 0 0 0

BTX Reformer 0 42 54 42 54 0 0 42 54 0 0

C4 Isomerization 15 50 0 27 27 -23 27 10 15 -40 14

C5/C6 Isomerization 157 157 300 157 129 0 -171 100 0 -57 -300

HF Alkylation 156 121 78 78 92 -44 14 78 78 -44 0

H2SO4 Alkylation 81 81 71 79 81 -2 10 40 65 -40 -6

Dimersol 3 0 0 0 0 0 0 0 0 0 0

Cat Poly 4 0 29 0 3 0 -26 0 8 0 -21

Isooctane 0 0 0 0 0 0 0 0 0 0 0

DHT - Total 1,056 827 644 874 610 47 -34 762 581 -66 -63

DHT 2nd RCT - Total 0 0 0 0 0 0 0 0 0 0 0

DHT Arom Saturation 6 0 0 0 0 0 0 0 0 0 0

NHT - Total Fd 1,038 621 1,007 597 966 -24 -41 495 924 -125 -83

CGH - Generic 282 242 233 1 183 -241 -49 0 179 -242 -54

CGH - Olefin Sat'n 282 0 0 0 0 0 0 0 0 0 0

FCCU Fd HDT 521 521 521 650 650 129 129 521 521 0 0

LSR Splitter 0 100 100 93 100 -7 0 77 71 -23 -29

LSR Bz Saturator 0 0 0 0 0 0 0 0 0 0 0

Reformate Saturator 0 2 10 1 1 -1 -9 1 1 -1 -9

Reformate Splitter 0 5 30 3 3 -2 -27 3 3 -2 -27

SDA 0 36 36 36 36 0 0 36 36 0 0

MTBE 0 0 0 0 0 0 0 0 0 0 0

TAME 0 0 0 0 0 0 0 0 0 0 0

Hydrogen Plant - Total MSCF 799,200 1,149 1,149 1,197 1,224 48 75 1,086 1,086 -63 -63

Lube Unit 66 24 24 24 24 0 0 24 24 0 0

Sulfur Plant 1,434 4 5 4 5 0 0 3 5 -1 0

Merox Jet 0 0 0 0 0 0 0 0 0 0 0

Merox Diesel 0 0 0 0 0 0 0 0 0 0 0

BTX Reformer - Tower feed 18 16 18 16 18 0 0 16 18 0 0

BTX Reformer - Extract feed 34 34 34 34 34 0 0 34 34 0 0

Toluene Dealkyation 0 0 0 0 0 0 0 0 0 0 0

Cumene 0 0 0 0 0 0 0 0 0 0 0

Cyclohexane 0 0 0 0 0 0 0 0 0 0 0

Reference 2005 in

2030 

100% E10 Case in

2030

E10 Case Relative to 

Ref Case

100% E15 Case in

2030

E15 Case Relative to

Ref Case

Actual 

Capacity

in 2009
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-80

Table 5-54 PADD 3 Unit Capacity and Throughput Volumes from LP Refinery Modeling 

(Thousand bbl/day)

PADD 3

Refinery Units Summer Winter Summer Winter Summer Winter Summer Winter Summer Winter

Crude Tower 8,735 7,741 7,111 7,477 7,036 -264 -75 7,203 6,648 -538 -464

Vacuum Tower 3,963 3,514 3,210 3,381 3,264 -133 54 3,275 3,022 -239 -188

Sats Gas Plant 0 635 427 624 471 -11 44 612 423 -23 -4

Unsats Gas Plant 0 679 675 690 612 10 -63 607 571 -73 -104

FCC DeC5 Tower 0 0 0 0 0 0 0 0 0 0 0

FCC 2,808 2,075 2,154 2,103 1,901 28 -252 1,820 1,794 -255 -360

FCC Splitter 0 6 1,153 161 435 155 -719 0 848 -6 -305

Hydrocracker 667 1,797 1,797 2,020 2,020 223 223 1,937 1,937 140 140

H-Oil Unit 0 0 0 0 0 0 0 0 0 0 0

Delayed Coker 1,350 992 853 933 881 -59 28 886 769 -106 -85

Visbreaker 51 0 0 0 0 0 0 0 0 0 0

Thermal Naphtha Splitter 0 120 103 113 107 -7 3 107 93 -13 -10

CRU Reformer 1,765 1,517 1,140 1,130 904 -387 -236 1,149 863 -368 -277

SRU Reformer 0 0 0 0 0 0 0 0 0 0 0

BTX Reformer 0 265 261 261 261 -4 0 261 261 -4 1

C4 Isomerization 50 50 0 50 0 0 0 50 0 0 0

C5/C6 Isomerization 199 130 58 229 50 99 -7 134 0 4 -58

HF Alkylation 295 148 148 148 148 0 0 148 148 0 0

H2SO4 Alkylation 290 234 235 252 197 18 -38 197 179 -37 -56

Dimersol 19 0 0 0 0 0 0 0 0 0 0

Cat Poly 23 0 0 0 0 0 0 0 0 0 0

Isooctane 0 0 0 0 0 0 0 0 0 0 0

DHT - Total 2,606 2,159 1,867 1,726 1,689 -434 -178 1,872 1,675 -288 -192

DHT 2nd RCT - Total 0 0 0 0 0 0 0 0 0 0 0

DHT Arom Saturation 4 0 0 0 0 0 0 0 0 0 0

NHT - Total Fd 2,284 1,648 1,198 1,474 1,076 -174 -122 1,828 1,300 180 102

CGH - Generic 905 1 1 26 1 25 0 1 1 0 0

CGH - Olefin Sat'n 905 0 0 0 0 0 0 0 0 0 0

FCCU Fd HDT 1,375 1,844 1,844 1,850 1,630 6 -213 1,562 1,559 -282 -285

LSR Splitter 0 1 1 100 1 99 0 1 1 0 0

LSR Bz Saturator 0 0 0 0 0 0 0 0 0 0 0

Reformate Saturator 0 327 76 91 10 -236 -65 330 272 3 196

Reformate Splitter 0 980 227 273 31 -707 -196 989 815 9 588

SDA 0 263 263 263 263 0 0 263 263 0 0

MTBE 28 0 0 0 0 0 0 0 0 0 0

TAME 3 0 0 0 0 0 0 0 0 0 0

Hydrogen Plant - Total MSCF 1,569,400 5,529 5,529 5,931 5,931 402 402 5,670 5,670 141 141

Lube Unit 1,169 440 440 440 440 0 0 440 440 0 0

Sulfur Plant 1,434 15 13 14 14 -1 1 13 12 -1 0

Merox Jet 0 0 0 0 0 0 0 0 0 0 0

Merox Diesel 0 0 0 0 0 0 0 0 0 0 0

BTX Reformer - Tower feed 0 96 94 94 94 -2 0 94 94 -2 0

BTX Reformer - Extract feed 0 211 208 208 208 -3 0 208 208 -3 0

Toluene Dealkyation 0 0 0 0 0 0 0 0 0 0 0

Cumene 0 0 0 0 0 0 0 0 0 0 0

Cyclohexane 0 0 0 0 0 0 0 0 0 0 0

Reference 2005 in

2030 

100% E10 Case in

2030

E10 Case Relative to 

Ref Case

100% E15 Case in

2030

E15 Case Relative to

Ref Case

Actual 

Capacity

in 2009
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-81

Table 5-55 PADD 4 and 5OC Unit Capacity and Throughput Volumes from LP Refinery 

Modeling (Thousand bbl/day)

PADD 4 and 5OC

Refinery Units Summer Winter Summer Winter Summer Winter Summer Winter Summer Winter

Crude Tower 1,796 1,513 1,547 1,418 1,442 -95 -104 1,400 1,400 -113 -146

Vacuum Tower 599 655 632 607 578 -48 -55 588 560 -67 -73

Sats Gas Plant 0 92 57 82 53 -10 -4 75 39 -17 -18

Unsats Gas Plant 0 113 115 104 105 -9 -10 101 100 -12 -14

FCC DeC5 Tower 0 0 0 0 0 0 0 0 0 0 0

FCC 344 392 382 360 345 -32 -37 344 331 -47 -51

FCC Splitter 0 218 222 202 204 -15 -18 182 196 -35 -26

Hydrocracker 91 141 191 141 191 0 0 141 191 0 0

H-Oil Unit 0 0 0 0 0 0 0 0 0 0 0

Delayed Coker 143 162 162 137 137 -25 -25 133 129 -29 -33

Visbreaker 56 0 0 7 3 7 3 26 8 26 8

Thermal Naphtha Splitter 0 19 19 16 16 -3 -3 17 15 -2 -3

CRU Reformer 0 0 0 0 0 0 0 0 0 0 0

SRU Reformer 287 129 201 95 73 -33 -128 69 27 -60 -174

BTX Reformer 0 0 0 0 0 0 0 0 0 0 0

C4 Isomerization 19 45 0 40 8 -5 8 40 6 -5 5

C5/C6 Isomerization 43 78 2 0 0 -78 -2 3 0 -75 -2

HF Alkylation 37 37 37 37 37 0 0 37 37 0 0

H2SO4 Alkylation 38 74 74 64 64 -10 -10 60 60 -14 -14

Dimersol 1 0 1 1 1 0 0 1 1 1 0

Cat Poly 10 0 0 0 0 0 0 0 0 0 0

Isooctane 0 0 0 0 0 0 0 0 0 0 0

DHT - Total 404 575 578 528 527 -48 -51 518 507 -57 -71

DHT 2nd RCT - Total 0 0 0 0 0 0 0 0 0 0 0

DHT Arom Saturation 0 0 0 0 0 0 0 0 0 0 0

NHT - Total Fd 326 313 349 292 297 -21 -53 288 287 -25 -62

CGH - Generic 66 78 74 66 60 -11 -14 65 60 -12 -14

CGH - Olefin Sat'n 66 0 0 0 0 0 0 0 0 0 0

FCCU Fd HDT 89 139 189 139 189 0 0 139 177 0 -12

LSR Splitter 0 100 100 89 88 -11 -12 85 100 -15 0

LSR Bz Saturator 0 0 0 0 0 0 0 0 0 0 0

Reformate Saturator 0 2 10 1 1 -1 -9 1 1 -1 -9

Reformate Splitter 0 5 30 3 3 -2 -27 3 3 -2 -27

SDA 0 28 28 28 28 0 0 28 28 0 0

MTBE 0 0 0 0 0 0 0 0 0 0 0

TAME 0 0 0 0 0 0 0 0 0 0 0

Hydrogen Plant - Total MSCF 373,900 705 785 679 826 -26 40 688 809 -18 23

Lube Unit 0 0 0 0 0 0 0 0 0 0 0

Sulfur Plant 1,434 2 2 2 2 0 0 2 2 0 0

Merox Jet 0 0 0 0 0 0 0 0 0 0 0

Merox Diesel 0 0 0 0 0 0 0 0 0 0 0

BTX Reformer - Tower feed 0 0 0 0 0 0 0 0 0 0 0

BTX Reformer - Extract feed 0 0 0 0 0 0 0 0 0 0 0

Toluene Dealkyation 0 0 0 0 0 0 0 0 0 0 0

Cumene 0 0 0 0 0 0 0 0 0 0 0

Cyclohexane 0 0 0 0 0 0 0 0 0 0 0

Reference 2005 in

2030 

100% E10 Case in

2030

E10 Case Relative to 

Ref Case

100% E15 Case in

2030

E15 Case Relative to

Ref Case

Actual 

Capacity

in 2009
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-82

Table 5-56 U.S. (except CA) Unit Capacity and Throughput Volumes from LP Refinery 

Modeling (Thousand bbl/day)

PADDs 1 -4 and 5OC

Refinery Unit Summer Winter Summer Winter Summer Winter Summer Winter Summer Winter

Crude Tower 15,403 13,049 14,515 12,146 13,717 -903 -798 11,567 13,145 -1,482 -1,370

Vacuum Tower 6,505 5,768 6,350 5,346 6,104 -422 -246 5,100 5,788 -668 -562

Sats Gas Plant 0 993 683 912 718 -81 35 865 629 -128 -54

Unsats Gas Plant 0 1,225 1,348 1,153 1,227 -72 -122 1,030 1,163 -195 -185

FCC DeC5 Tower 0 0 0 0 0 0 0 0 0 0 0

FCC 4,786 4,008 4,506 3,694 4,018 -314 -489 3,298 3,846 -710 -660

FCC Splitter 0 1,114 2,494 694 1,658 -420 -837 506 2,030 -608 -464

Hydrocracker 1,013 2,223 2,303 2,446 2,526 223 223 2,363 2,443 140 140

H-Oil Unit 0 0 0 0 0 0 0 0 0 0 0

Delayed Coker 1,912 1,329 1,505 1,222 1,475 -108 -30 1,127 1,334 -202 -171

Visbreaker 107 0 0 7 3 7 3 26 8 26 8

Thermal Naphtha Splitter 0 159 181 147 178 -12 -3 137 161 -23 -20

CRU Reformer 2,794 2,048 1,976 1,596 1,572 -453 -404 1,613 1,513 -436 -463

SRU Reformer 287 129 201 95 73 -33 -128 69 27 -60 -174

BTX Reformer 0 344 351 340 351 -4 0 340 351 -4 1

C4 Isomerization 116 199 0 149 35 -49 34 132 20 -67 20

C5/C6 Isomerization 406 415 455 418 212 2 -244 253 8 -162 -447

HF Alkylation 528 346 283 302 297 -44 14 302 282 -44 0

H2SO4 Alkylation 438 513 448 488 363 -25 -85 374 324 -139 -124

Dimersol 23 0 1 1 1 0 0 1 1 1 0

Cat Poly 42 1 60 1 35 0 -25 8 40 7 -20

Isooctane 0 0 0 0 0 0 0 0 0 0 0

DHT - Total 4,400 3,993 3,831 3,428 3,398 -565 -432 3,462 3,336 -531 -495

DHT 2nd RCT - Total 0 0 0 0 0 0 0 0 0 0 0

DHT Arom Saturation 10 0 0 0 0 0 0 0 0 0 0

NHT - Total Fd 3,921 2,844 2,952 2,527 2,625 -317 -327 2,763 2,797 -81 -155

CGH - Generic 1,373 481 466 252 402 -229 -64 225 398 -257 -68

CGH - Olefin Sat'n 1,373 0 0 0 0 0 0 0 0 0 0

FCCU Fd HDT 2,029 2,964 3,014 3,035 2,866 71 -148 2,619 2,653 -345 -360

LSR Splitter 0 202 252 283 213 81 -39 164 212 -38 -40

LSR Bz Saturator 0 0 0 0 0 0 0 0 0 0 0

Reformate Saturator 0 351 97 112 13 -239 -83 342 275 -10 178

Reformate Splitter 0 1,054 290 337 40 -717 -250 1,025 824 -29 534

SDA 0 359 359 359 359 0 0 359 359 0 0

MTBE 31 0 0 0 0 0 0 0 0 0 0

TAME 3 0 0 0 0 0 0 0 0 0 0

Hydrogen Plant - Total MSCF 1,238,369 8,029 8,399 8,346 8,769 317 370 7,962 8,360 -67 -38

Lube Unit 1,380 516 516 516 516 0 0 516 516 0 0

Sulfur Plant 5,736 21 21 20 22 -1 0 19 20 -2 -1

Merox Jet 0 0 0 0 0 0 0 0 0 0 0

Merox Diesel 0 0 0 0 0 0 0 0 0 0 0

BTX Reformer - Tower feed 127 124 125 123 125 -2 0 123 125 -2 0

BTX Reformer - Extract feed 268 268 265 265 265 -3 0 265 265 -3 0

Toluene Dealkyation 0 0 0 0 0 0 0 0 0 0 0

Cumene 0 0 0 0 0 0 0 0 0 0 0

Cyclohexane 0 0 0 0 0 0 0 0 0 0 0

100% E15 Case in

2030

E10 Case Relative to 

Ref Case

100% E10 Case in

2030

Reference 2005 in

2030 

E15 Case Relative to

Ref Case

Actual 

Capacity

in 2009
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-83

Table 5-57 PADD 1 Gasoline Qualities for the E10 and E15 Cases relative to the Reference Case

CG RFG Pool CG RFG Pool CG RFG Pool CG RFG Pool CG RFG Pool CG RFG Pool

Energy (MMBTU/bbl) 5.18 5.09 5.14 5.11 4.93 5.02 5.05 5.04 5.04 5.02 4.91 4.98 4.93 4.93 4.93 4.87 4.82 4.85

Density (lb/bbl) 257.70 259.91 258.80 255.81 252.05 253.91 257.59 260.94 258.89 255.60 255.72 255.65 259.60 260.61 260.01 257.91 255.72 257.01

Sulfur (ppm) 27.5 23.3 25.4 22.4 16.9 19.6 23.7 23.6 23.7 20.9 18.0 19.8 28.0 28.0 28.0 12.6 17.6 14.7

% at 200 45.8 50.8 48.3 44.6 62.4 53.6 54.8 51.1 53.3 53.3 59.4 55.7 59.8 58.7 59.4 60.9 66.7 63.3

% at 300 76.6 81.4 79.0 75.8 89.2 82.6 81.6 81.2 81.4 78.1 84.6 80.7 81.5 84.1 82.6 83.0 89.8 85.8

RVP (psi) 7.92 6.54 7.23 12.61 13.19 12.90 9.27 7.00 8.39 13.78 13.39 13.62 8.27 7.00 7.75 13.29 12.92 13.14

T10 (F) 131.2 140.0 135.6 101.4 97.7 99.6 122.6 137.1 128.2 94.0 96.5 95.0 129.0 137.1 132.3 97.1 99.5 98.1

T50 (F) 208.5 198.2 203.3 210.9 174.5 192.5 190.1 197.5 193.0 193.0 180.5 188.1 179.8 182.1 180.7 177.5 165.7 172.7

T90 (F) 358.5 336.6 347.6 362.1 301.3 331.4 336.0 337.6 336.7 351.6 322.1 340.0 336.3 324.3 331.4 329.5 298.3 316.7

Driveability 1180.7 1141.2 1161.0 1147.0 971.4 1058.2 1090.3 1135.7 1107.9 1071.7 1008.4 1046.8 1069.1 1076.0 1072.0 1007.7 944.7 981.9

Vapor Lock 143.8 149.5 146.6 119.1 109.7 114.3 133.4 146.9 138.7 109.7 109.7 109.7 137.1 144.3 140.0 109.7 109.7 109.7

Aromatics (vol%) 27.9 20.2 24.1 27.0 14.8 20.8 22.0 20.0 21.2 20.3 18.5 19.6 18.5 17.6 18.1 18.8 14.0 16.8

Benzene (vol%) 0.58 0.57 0.58 0.61 0.56 0.59 0.51 0.57 0.53 0.52 0.58 0.54 0.73 0.89 0.79 0.46 0.55 0.50

Olefins (vol%) 11.0 10.4 10.7 10.6 16.8 13.8 9.1 10.9 9.8 10.2 16.0 12.5 8.5 9.8 9.0 7.9 17.9 12.0

Alcohol (vol%) 0.50 4.89 2.68 0.00 4.94 2.50 10.00 10.00 10.00 10.00 10.00 10.00 15.00 15.00 15.00 15.00 15.00 15.00

Oxygen (wt%) 0.74 2.56 1.65 0.15 2.50 1.34 3.74 3.69 3.72 3.77 3.77 3.77 5.57 5.55 5.56 5.61 5.65 5.63

Volumes (kbbl/day) 1,399,252 1,387,051 2,786,303 1,338,953 1,368,062 2,707,015 1,760,831 1,121,463 2,882,294 1,698,176 1,106,878 2,805,054 1,736,796 1,204,323 2,941,119 1,619,601 1,125,767 2,745,368

Gasoline Qualities and 

Volume for PADD 1

2005 in 2030 E10 E15

Summer Winter Summer Winter Summer Winter

Table 5-58 PADD 2 Gasoline Qualities for the E10 and E15 Cases relative to the Reference Case

CG RFG Pool CG RFG Pool CG RFG Pool CG RFG Pool CG RFG Pool CG RFG Pool

Energy (MMBTU/bbl) 5.18 5.02 4.97 4.99 4.75 4.95 4.99 5.03 4.99 4.87 4.89 4.88 4.94 4.93 4.94 4.83 4.81 4.83

Density (lb/bbl) 257.79 258.55 259.05 255.20 251.15 254.54 258.09 259.13 258.25 256.71 256.01 256.59 260.15 259.67 260.07 258.46 256.91 258.20

Sulfur (ppm) 27.4 9.7 20.7 25.0 16.9 23.7 23.3 17.9 22.4 24.1 24.1 24.1 28.0 28.0 28.0 26.8 26.7 26.8

% at 200 45.7 53.7 56.1 48.8 68.4 51.9 53.4 49.7 52.8 58.9 59.0 58.9 57.5 55.8 57.2 62.7 64.8 63.0

% at 300 76.6 85.0 84.6 79.7 93.7 82.0 81.7 82.2 81.7 83.9 84.9 84.1 80.3 81.3 80.5 84.2 86.7 84.6

RVP (psi) 7.91 7.00 8.73 13.97 12.70 13.76 9.28 7.00 8.92 13.31 13.31 13.31 8.28 7.00 8.08 13.07 12.93 13.05

T10 (F) 131.2 137.1 126.0 92.8 100.9 94.1 122.6 137.1 124.9 97.0 97.0 97.0 128.9 137.1 130.2 98.5 99.4 98.6

T50 (F) 208.6 192.3 187.4 202.3 162.3 195.8 192.9 200.4 194.1 181.7 181.5 181.6 184.6 188.0 185.1 173.9 169.6 173.2

T90 (F) 358.7 320.2 322.3 344.5 280.8 334.2 335.6 333.0 335.2 325.4 320.8 324.6 341.6 337.2 340.9 323.8 312.8 322.0

Driveability 1181.5 1102.5 1073.5 1090.7 918.9 1062.8 1098.0 1139.7 1104.7 1015.8 1010.7 1015.0 1088.7 1106.7 1091.6 993.4 970.6 989.7

Vapor Lock 143.9 146.0 135.9 110.3 110.3 110.3 133.9 147.4 136.0 110.3 110.3 110.3 137.8 145.3 139.0 110.3 110.3 110.3

Aromatics (vol%) 27.9 18.9 22.2 24.7 12.3 22.7 22.5 19.5 22.0 20.4 18.2 20.0 20.3 19.2 20.2 20.0 16.9 19.5

Benzene (vol%) 0.58 0.56 0.54 0.58 0.57 0.58 0.51 0.56 0.52 0.50 0.55 0.51 0.55 0.76 0.58 0.67 0.65 0.67

Olefins (vol%) 11.5 3.8 8.8 13.0 6.8 12.0 8.8 6.7 8.5 10.7 11.2 10.7 7.3 7.0 7.2 10.5 10.6 10.6

Alcohol (vol%) 0.42 10.00 8.80 0.00 10.00 1.62 10.00 10.00 10.00 10.00 10.00 10.00 15.00 15.00 15.00 15.00 15.00 15.00

Oxygen (wt%) 0.40 3.73 3.27 0.00 3.84 0.62 3.74 3.72 3.73 3.76 3.77 3.76 5.56 5.57 5.56 5.60 5.63 5.60

Volumes (kbbl/day) 1,772,351 329,642 1,759,558 1,685,250 326,672 2,011,922 1,802,319 342,519 2,144,838 1,751,081 339,433 2,090,514 1,841,582 342,519 2,184,101 1,751,081 339,433 2,090,514

Gasoline Qualities and 

Volume for PADD 2

Summer Winter Summer Winter Summer Winter

2005 in 2030 E10 E15
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-84

Table 5-59 PADD 3 Gasoline Qualities for the E10 and E15 Cases relative to the Reference Case

CG RFG Pool CG RFG Pool CG RFG Pool CG RFG Pool CG RFG Pool CG RFG Pool

Energy (MMBTU/bbl) 5.18 5.16 5.18 5.13 5.18 5.14 5.07 5.11 5.08 5.07 5.00 5.05 4.95 4.95 4.95 4.91 4.84 4.89

Density (lb/bbl) 258.31 258.54 258.37 256.15 254.74 255.78 258.18 259.96 258.66 256.90 255.26 256.48 260.34 260.35 260.35 258.80 254.37 257.67

Sulfur (ppm) 27.6 24.8 26.8 25.7 22.4 24.7 24.1 24.9 24.3 22.8 23.1 22.9 27.8 26.0 27.3 13.1 21.6 15.3

% at 200 44.8 51.4 46.6 43.7 56.5 45.8 54.0 49.0 52.7 50.7 53.7 51.4 58.8 58.3 58.6 59.9 66.5 61.6

% at 300 75.9 82.7 77.8 75.5 83.3 76.8 81.1 80.8 81.0 76.8 79.8 77.6 80.6 83.1 81.3 83.4 90.9 85.3

RVP (psi) 7.94 6.54 7.55 12.47 11.65 12.34 8.85 7.00 8.35 12.02 11.83 11.97 7.85 7.00 7.61 11.43 11.00 11.32

T10 (F) 131.1 140.0 133.5 102.3 107.5 103.2 125.3 137.1 128.5 105.2 106.4 105.5 131.7 137.1 133.1 108.9 111.6 109.6

T50 (F) 210.4 197.0 206.7 212.6 186.6 208.4 191.6 201.8 194.4 198.4 192.3 196.9 181.9 182.8 182.2 179.7 166.2 176.2

T90 (F) 361.9 330.7 353.3 363.6 328.2 357.6 338.1 339.4 338.4 357.6 343.9 354.1 340.3 328.9 337.2 327.5 293.7 318.9

Driveability 1189.8 1131.7 1173.7 1154.8 1049.1 1137.5 1100.8 1150.5 1114.2 1110.6 1080.5 1102.9 1083.6 1082.8 1083.4 1029.9 959.7 1012.0

Vapor Lock 144.0 149.3 145.5 120.1 120.1 120.1 135.9 147.6 139.1 120.1 120.1 120.1 139.7 144.4 141.0 120.1 120.1 120.1

Aromatics (vol%) 29.2 19.7 26.6 27.6 20.0 26.1 22.4 19.5 21.6 20.8 18.4 20.2 19.1 15.6 18.1 17.9 11.0 16.2

Benzene (vol%) 0.59 0.58 0.59 0.59 0.58 0.58 0.52 0.51 0.52 0.50 0.57 0.52 0.73 0.90 0.77 0.40 0.59 0.45

Olefins (vol%) 11.5 1.1 8.6 11.5 9.8 11.0 9.4 2.7 7.6 8.4 8.5 8.5 7.9 2.8 6.5 7.5 13.5 9.0

Alcohol (vol%) 0.90 5.00 2.03 0.00 4.00 0.63 10.00 10.00 10.00 10.00 10.00 10.00 15.00 15.00 15.00 15.00 15.00 15.00

Oxygen (wt%) 0.67 2.54 1.19 0.00 2.50 0.40 3.73 3.71 3.73 3.75 3.78 3.76 5.55 5.55 5.55 5.59 5.68 5.61

Volumes (kbbl/day) 1,257,764 480,771 1,738,535 1,400,629 489,720 1,890,349 1,306,897 479,961 1,786,858 1,492,196 508,849 2,001,045 1,306,897 499,552 1,806,449 1,433,678 488,895 1,922,573

Gasoline Qualities and 

Volume for PADD 3

Summer Winter

2005 in 2030 E10 E15

Summer Winter Summer Winter
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-85

Table 5-60 PADD 4 and 5OC Gasoline Qualities for the E10 and E15 Cases relative to the Reference Case

CG RFG Pool CG RFG Pool CG RFG Pool CG RFG Pool CG RFG Pool CG RFG Pool

Energy (MMBTU/bbl) 5.06 0.00 5.06 5.00 0.00 5.00 4.95 0.00 4.95 4.89 0.00 4.89 4.89 0.00 4.89 4.80 0.00 4.80

Density (lb/bbl) 258.41 0.00 258.41 253.46 0.00 253.46 258.55 0.00 258.55 254.99 0.00 254.99 260.44 0.00 260.44 256.45 0.00 256.45

Sulfur (ppm) 24.8 0.0 24.8 24.4 0.0 24.4 28.0 0.0 28.0 28.0 0.0 28.0 28.0 0.0 28.0 25.8 0.0 25.8

% at 200 53.3 0.0 53.3 59.2 0.0 59.2 62.1 0.0 62.1 64.4 0.0 64.4 64.6 0.0 64.6 70.1 0.0 70.1

% at 300 90.7 0.0 90.7 92.2 0.0 92.2 91.6 0.0 91.6 92.1 0.0 92.1 92.2 0.0 92.2 96.0 0.0 96.0

RVP (psi) 8.07 0.00 8.07 13.00 0.00 13.00 9.08 0.00 9.08 12.96 0.00 12.96 8.08 0.00 8.08 12.59 0.00 12.59

T10 (F) 130.3 0.0 130.3 99.0 0.0 99.0 123.8 0.0 123.8 99.2 0.0 99.2 130.2 0.0 130.2 101.5 0.0 101.5

T50 (F) 193.0 0.0 193.0 181.1 0.0 181.1 175.1 0.0 175.1 170.5 0.0 170.5 170.0 0.0 170.0 158.9 0.0 158.9

T90 (F) 294.6 0.0 294.6 287.8 0.0 287.8 290.3 0.0 290.3 288.3 0.0 288.3 287.5 0.0 287.5 270.5 0.0 270.5

Driveability 1069.1 0.0 1069.1 979.5 0.0 979.5 1001.4 0.0 1001.4 948.5 0.0 948.5 992.8 0.0 992.8 899.4 0.0 899.4

Vapor Lock 140.4 0.0 140.4 111.9 0.0 111.9 131.9 0.0 131.9 110.3 0.0 110.3 136.4 0.0 136.4 110.3 0.0 110.3

Aromatics (vol%) 20.4 0.0 20.4 20.3 0.0 20.3 17.4 0.0 17.4 15.3 0.0 15.3 15.7 0.0 15.7 12.7 0.0 12.7

Benzene (vol%) 1.17 0.00 1.17 1.12 0.00 1.12 1.19 0.00 1.19 1.18 0.00 1.18 1.19 0.00 1.19 0.73 0.00 0.73

Olefins (vol%) 9.7 0.0 9.7 8.8 0.0 8.8 8.1 0.0 8.1 7.8 0.0 7.8 7.8 0.0 7.8 7.3 0.0 7.3

Alcohol (vol%) 0.00 0.00 0.00 0.00 0.00 0.00 10.00 0.00 10.00 10.00 0.00 10.00 15.00 0.00 15.00 15.00 0.00 15.00

Oxygen (wt%) 0.01 0.00 0.01 0.00 0.00 0.00 3.73 0.00 3.73 3.78 0.00 3.78 5.55 0.00 5.55 5.64 0.00 5.64

Volumes (kbbl/day) 889,157 0 889,157 733,803 0 733,803 718,402 0 718,402 762,467 0 762,467 718,402 0 718,402 762,467 0 762,467

Gasoline Qualities and 

Volume for PADDs 4 & 

5OC

2005 in 2030 E10 E15

Summer Winter Summer Winter Summer Winter
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-86

Table 5-61 U.S. (except CA) Gasoline Qualities for the E10 and E15 Cases relative to the Reference Case

CG RFG Pool CG RFG Pool CG RFG Pool CG RFG Pool CG RFG Pool CG RFG Pool

Energy (MMBTU/bbl) 5.16 5.10 5.10 5.06 4.96 5.03 5.02 5.06 5.03 4.97 4.93 4.96 4.93 4.93 4.93 4.86 4.82 4.85

Density (lb/bbl) 257.99 259.40 258.71 255.37 252.52 254.52 258.01 260.38 258.62 256.20 255.65 256.06 260.06 260.39 260.15 258.11 255.59 257.45

Sulfur (ppm) 27.1 21.6 24.5 24.4 18.1 22.5 24.2 22.9 23.9 23.3 20.4 22.6 28.0 27.5 27.8 19.0 20.2 19.3

% at 200 46.8 51.4 50.4 47.8 62.0 51.7 55.1 50.4 53.9 55.8 57.9 56.3 59.4 58.1 59.1 62.5 66.3 63.5

% at 300 78.8 82.2 81.5 79.3 88.5 81.9 82.8 81.3 82.4 81.4 83.4 81.9 82.3 83.4 82.6 85.3 89.5 86.4

RVP (psi) 7.95 6.61 7.78 13.07 12.77 13.00 9.15 7.00 8.60 13.07 12.97 13.04 8.15 7.00 7.84 12.65 12.44 12.59

T10 (F) 131.0 139.5 132.1 98.5 100.4 98.9 123.4 137.1 126.9 98.5 99.2 98.7 129.7 137.1 131.7 101.2 102.5 101.5

T50 (F) 206.4 197.0 199.0 204.3 175.4 196.4 189.4 199.1 191.9 187.9 183.8 186.9 180.6 183.2 181.3 174.4 166.5 172.3

T90 (F) 348.7 332.9 336.2 346.2 304.3 334.5 330.5 337.3 332.2 336.7 327.5 334.3 332.7 327.6 331.3 319.1 299.7 314.1

Driveability 1164.5 1133.3 1131.2 1106.9 981.0 1072.0 1083.8 1140.1 1098.3 1048.3 1027.6 1043.0 1069.2 1082.8 1072.8 994.1 953.0 983.4

Vapor Lock 143.3 148.9 142.9 115.4 112.1 114.5 134.0 147.2 137.4 112.7 112.5 112.6 137.8 144.5 139.6 112.6 112.4 112.6

Aromatics (vol%) 27.0 19.9 23.8 25.5 15.6 22.6 21.7 19.8 21.2 19.8 18.4 19.4 18.9 17.4 18.5 18.1 13.8 17.0

Benzene (vol%) 0.68 0.57 0.64 0.67 0.56 0.64 0.60 0.55 0.59 0.60 0.57 0.59 0.73 0.87 0.77 0.55 0.58 0.56

Olefins (vol%) 11.1 7.4 9.6 11.4 13.7 12.1 8.9 8.1 8.7 9.6 13.2 10.5 7.9 7.6 7.8 8.5 15.5 10.4

Alcohol (vol%) 0.48 5.68 3.69 0.00 5.49 1.53 10.00 10.00 10.00 10.00 10.00 10.00 15.00 15.00 15.00 15.00 15.00 15.00

Oxygen (wt%) 0.49 2.73 1.73 0.04 2.70 0.77 3.74 3.70 3.73 3.76 3.77 3.76 5.56 5.55 5.56 5.60 5.66 5.62

Volumes (kbbl/day) 5,318,524 2,197,464 7,173,553 5,158,635 2,184,454 7,343,089 5,588,449 1,943,943 7,532,392 5,703,920 1,955,160 7,659,080 5,603,677 2,046,394 7,650,071 5,566,827 1,954,095 7,520,922

Gasoline Qualities and 

Volume for PADDs 1 -

4 & 5OC

2005 in 2030 E10 E15

Summer Winter Summer Winter Summer Winter
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

5-87

References

1 Annual Energy Outlook 2006, Energy Information Administration, Department of Energy

2 Worldwide Report, Oil and Gas Journal. www.ogjonline.com, December 22, 2005.

3

Petroleum Supply Annual 2004, Volume 1, Table 38. Energy Information Administration, Department of Energy.

4

Petroleum Supply Annual 2004, Volume 1, Tables 22 – 25. Energy Information Administration, Department of 

Energy.

5

Petroleum Supply Annual 2005, Volume 1, Table 17. Energy Information Administration, Department of Energy.

6

Petroleum Refining Technology and Economics 4th Edition. James Gary and Glenn Handerwork. Page 65, Table 

4.7

7 The Challenges & Opportunities of 10 ppm Sulfur Gasoline. Jay Ross, Delphine Largeteau, Marc Laborde, Larry 

Wisdom. NPRA Annual Meeting, AM-11-57, March, 2011, page 6

8 The Benefits of Cat Feed Hydrotreating and the Impact of Feed Nitrogen on Catalyst Stability. Brian Moyse at 

Haldor Topso. NPRA Annual Meeting, AM-10-167, March 2010, Page 2

9 Handbook of Petroleum Refining Processes, Third Edition. Robert Meyers, page 11.39

10 Refining Economics of a National Low Sulfur, Low RVP Gasoline Standard, Performed for The International 

Council for Clean Transportation by Mathpro, October 25, 2011.

11 Refining Economics of a National Clean Gasoline Standard for PADDs 1-3; for The Alliance of Automobile 

Manufacturers by Mathpro, June 27, 2008. 

12 Potential Supply and Cost Impacts of Lower Sulfur, Lower RVP Gasoline; prepared for The American Petroleum 

Institute by Baker and O’Brien; July 2011.

13 Schink, George R., Singer, Hal J., Economic Analysis of the Implications of Implementing EPA’s Tier 3 Rules, 

prepared for the Emissions Control Technology Association, June 14, 2012.

14 An Assessment of the Impact of Nonroad Diesel Fuel Sulfur Regulation on Distillate Fuel Production and 

Availability in the U.S., prepared for the American Petroleum Institute by Baker and O’Brien, July 2003.

15 Petroleum Supply Annual 2003 – 2010; Table 24, Imports of Crude Oil and Petroleum Products, Energy 

Information Administration. 

16 Petroleum Supply Annual 2003 – 2010; Table 31, Exports of Crude Oil and Petroleum Products by PAD 

Destination.

17 Petroleum Supply Annual 2011, Refinery Capacity Report, Table 13 Refineries Permanently Shutdown by PAD 

District Between January 1, 1990 and January 1, 2010; Energy Information Administration.

18 Petroleum Supply Annual 2003 – 2011, Refinery Capacity Report, Table 11 New, Shutdown and Reactivated 

Refineries During 2010; Energy Information Administration.

19 Petroleum and Other Liquids; Number and Capacity of Petroleum Refineries, Data, Total Number of Operable 

Refineries and Total Number of Operating Refineries, 1990 – 2010; Energy Information Administration.
*** E.O. 12866 Review – Revised Version – Do Not Cite, Quote, or Release During Review ***

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20 Gary, J., & Handewerk, G. (2001). Petroleum Refining: Technology and Economics. (4th ed.). CRC Press. p 256.

21 Gary, J., & Handewerk, G. (2001). Petroleum Refining: Technology and Economics. (4th ed.). CRC Press. p 337.

22 Keller, P. (2012, January). New Source Review Permitting Impact Analysis for Proposed Tier 3 Gasoline 

Program. Memorandum to the docket.