From: Lester Wy borny To: Docket EPA

- HQ

- OAR

- 2011

- 0 135 Date: April 11 , 201 3 Subject: Refinery

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- Refinery Gasoline Sulfur Cost Model

- Response to Peer Review

Comments 1. Background EPA developed a cost model which estimates the cost of reducing gasoline sulfur levels on a refinery

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- refinery

basis for the Tier 3 notice of proposed rulemaking (NPRM) . To enhance the cost estimation ability of the cost model, EPA utilized detailed, refinery

- specific information in the cost model, much of which was confidential business. The confidential business information included gasoline volume and quality information that refiners report to EPA as part of the compliance requirements . To allow EPA to be tter estimate the sulfur levels of each refinery ’s gasolin e, EPA obtained crude sulfur level data for each refinery from the Energy In formation Administration (EIA). EPA also obtained actual refinery unit volumetric throughput information from EIA for the fluidized catal yt ic cracker (FCC unit) 1 , atmospheric crud e tower , hydrocrackers, and coker units. All of this data is considered confidential business information by EIA. Finall y, EPA obtained detailed FCC naphtha desulfurization information from several vendors. That information was also confidential business information

. We contracted RTI to facilitate the independent peer review s of the refinery

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- refinery cost model

. RTI sub contracted Charles Lieder as one peer reviewer (Reviewer 1) . Charles Lieder was the Technology Task Group Refining Committee Chair for the National Petroleum Council (NPC) when the NPC was assessing the cost of clean fuels, including low sulfur gasoline , prior to Tier 2. Complying with Tier 3 requires revamping the FCC postreaters installed for Tier 2 , or installing scaled down grassroots postreaters like those installed for Tier 2 , which is why

Charles Lieder’s experience with Tier 2, particularl y the Tier 2 postreating technologies, is valuable for this project

. The second and third peer reviewer s w ere Martin Tallet and Dan Dunbar of Ensy s (Reviewer s 2 and 3 )

. Ensy s is a refining industry contractor and uses its World Model, an LP refinery model, to conduct energy and cost analy ses for the refining industry , as well as the government

. The refinery - by

- refinery cost model that we are using for cost estimation is broadly similar to an LP refinery model like the World Model. Thus, Martin Tallet’s and Dan Dunbar’s experience running their World Model provides them good background experience for review ing the refinery

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- refinery cost model.

1 The FCC unit cracks the heavy, sour part of crude oil and produces a significant amount of the gasoline blendstock which makes up a refinery’s gas oline pool. Therefore, the FCC unit is responsible for over 90 percent of the sulfur which ends up in the gasoline pool for those refineries with an FCC unit

.Rather than deal with all the challenges and complexity associated with providing the contractors with access to all the CBI data discussed above, we discovered that we could remove the EPA fuel quality data from the cost model and still allow a peer review of the model

. The refinery

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- refinery

cost model then defaulted to estimating its own gasoline volume for each refinery based on the capacities of the refinery units and production volumes of straight run naphtha as estimated from the volume of crude oil that was refined. The peer reviews were conducted in earl y November and earl y December of 2011

. Draft reports were provided to EPA in December, and the completed peer review reports were provided to EPA in earl y J anuary of 2012

. The purpose of this document is to respond to the comments provided b y the peer reviewers. We do not believe the peer reviewers identified any problems with the draft refinery model that warranted changes prior to being used for the NPRM. However, the peer reviewers did provide valuable and meaningful review and comment, and we intend to take their comments into consideration in conducting the refinery modeling used to support the final rule. Other than committing to fixing min or calculation errors the peer reviewers discovered in the cost model, this document is not intended to summarize any specific changes that the refinery cost model will undergo for the final rule. We will assess whether we need to make changes to the cost model once we have received the public comments that are provided on the draft cost analy sis

. 2. Peer Review Comments and EPA Response A. Calculation Errors Comments As indicated in the comments primarily by Reviewer 1, but also by

Reviewer s 2 and 3 , the refinery

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- refinery

cost model contained a small number of errors which, when corrected, has a negligible impact on the overall gasoline desulfurization costs estimated by the model. The errors identified include:

 T he formula which estimates steam costs solel y picked up the variable costs , omitting the capital cost portion of the cost. Also, the capital cost portion of the equation applied a capital cost amortization factor which was too low (0.04 instead of 0.11). The peer reviewer estimated th at these errors resulted in underestimating the desulfurization cost by 0.02 cents per gallon.
 F or a subset group of refineries, the refinery model was not using the maximum FCC naphtha throughput volume as estimated using actual 2009 FCC unit throughput v olumes, which slightly

underestimates the capital costs for the postreater revamps for those refineries.
 There was an e rror in the desulfurization technology that was applied to one refin ery for both for the 10 and 5 ppm cost estimate

. F or PADD 4 refineries the model was referring to PADD 5 utility costs instead of those for PADD 4 , slightl y overestimating the desulfurization costs

.  There was an error in the a pplication of light straight run (L SR) treating costs for 1 refinery

. Response T hese calculation errors will be corrected for the final rule cost anal ys is. B. Capital cost estimates Comments Both reviewers commented on the capital cost estimate information provided by the vendor companies. Reviewer 1’s comments were two

- fold. First, the commenter suggested obtaining more information from the vendors that would allow us to better assess whether the capital cost estimates provided by the vendors are realistic. Second, the commenter suggested hiring a h ydrotreating capital cost expert t o review the cost information to determine if the capital cost information is realistic, and if not what adjustments should be made. Reviewer s 2 and 3 also had reservations about the capital costs that we used. One comment that Reviewer s 2 and 3 made wa s that often vendor cost estimates tend to be biased low and that they should be increased to account for this potential bias. To account for this potential bias, Reviewer s 2 and 3 suggested using a 20% contingency factor. Reviewer s 2 and 3 also commented that the capital cost information provided by one of the vendor s was internally inconsistent. This vendor had supplied different capital and operating cost estimates for revamping an existing FCC naphtha postreater based on whether the FCC naphtha exiting the FCC unit contains 2400, 800 and 200 ppm sulfur before being desulfurized in a Tier 2 gasoline desulfurization unit. Reviewer s 2 and 3 noted that a revamp to an FCC postreater treating FCC naphtha with 2 0 00 ppm sulfur was highest, but that the revamp to a FCC postreater treating FCC naphtha with 200 ppm sulfur was next highest in revamp costs, not an FCC naphtha postreater treating 800 ppm

– arguing that the revamped capital costs did not seem consistent across the different revamp scen arios. Reviewer s 2 and 3 also observed that the ratio of revamp capital costs to grassroots capital costs were not consistent with the ratio reported by Handework. In his book [ Petroleum Refining Technology and Economics, 4 th edition, Hand e werk ], Handework estimates that revamp costs for a unit are 30 percent of the capital cost for a grassroots unit

. Reviewers 2 and 3 also noted that we used the vendor

- provided capital cost for a grassroots hydrotreater for revamping an FCC postreater to desulfurize FCC naphtha from 75 ppm to 10 ppm , for processes to desulfurize FCC naphtha from 75 ppm to 25 ppm. ResponseIn an attempt to verify the level of capital cost provided by the vendors, w e performed two cross check s of the capital costs. For the grassroots FCC postreater, we compared our FCC postreater grassroots capital cost estimate to two other capital cost estimates to serve as a check on our costs. One crosscheck was the grassroots capital cost for Tier 3 used by Mathpro for an FCC postreater to reduce their gasoline’s sulfur content to 10 ppm from their current Tier 2 sulfur levels

. Mathpro estimated a capital cost of $1830/bbl/day , but Mathpro’s cost estimate includes outside battery

limits (OSBL) costs. Since the vendor 1 cost estimate of $1500/bbl/day only includes inside battery

limits costs (I SBL), to compare our costs with Mathpro’s we need ed to apply our IS BL to OSBL cost factor which is 1.25. When we appl y our OSBL cost factor to our grassroots FCC postr eater costs, vendor 1 costs increase to $1870/bbl/day , which is slightl y higher than Mathpro’s capital cost estimate , but because the values are so close, they are essentially consistent with each other

. For the second crosscheck, w e compare d our grassroots capital costs to the capital cost estimated by

Jacobs Engineering , which we have access to as part of the Haverly LP refinery model. The Jacobs Engineering data base is provided with the LP refinery cost model to enable the refining costs of us ing existing and new refinery units to be estimated . What is important about this is that Jacobs Engineering engineered Tier 2 FCC postreaters for refiners and therefore has first hand access to capital cost information that refiners faced when complying with the 30 ppm Tier 2 gasoline sulfur standard. Jacobs’ ISBL cost for an FCC postreater is $2200/bbl/day

. Jacob’s capital cost is higher than the $1500/bbl/day I SBL cost that we are using based on the vendor data , h owever, the desulfurization scenarios are quite different. While the grassroots hydrotreater designed for Tier 3 is for desulfurizing FCC naphtha from 100 ppm to 25 or 10 ppm, a FCC postreater designed for Tier 2 was ty pically designed to desulfurize FCC naphtha from an average of about 800 ppm down to 75 ppm. 2 Thus, the Tier 2 FCC postreater is required to remove about ten times more sulfur than the Tier 3 treater. H igher FCC naphtha sulfur levels tends to cause increased recombination reactions which limits the ability of a simple hydro treater to achieve the necessary

desulfurization duty

with reasonable levels of octane loss and hydrogen consumption. To avoid high octane loss and h ydrogen consumption, the FCC postreater vendors designed their postreaters with additional subunits to lim it the recombination reactions. The additional subunits varied depending on the vendor. For Scanfining, a Zeromer or Exomer unit could be added. For Selectfining, a second stage could be added. For Axens, a splitter and Selective Hydrogenation Unit (SHU) could be added. We asked the vendor wh ich supplied the capital cost information that we used for a Tier 3 FCC postreater what the capital costs would be if the unit were designed to desulfurize FCC naphtha with a significantl y higher sulfur level (i.e. , 800 ppm) for Tier 2 rather than Tier 3 . The vendor responded that one or two sub unit s would likely be added as well as using a larger reactor. The vendor estimated the average capital cost for these additions or changes would be about $500/bbl/day , thus increasing the vendor cost for a

2 US Petroleum Refining Assuring the Adequacy and Affordability of Cleaner Fuels, National Pe troleum Council, June 2000.grassroots FCC postreater designed to treat an 800 ppm FCC naphtha feed to about

$2000/bbl/day , within 10 percent of the Jacobs capital cost estimate. We will further evaluate the capital costs for a grassroots treater to determine whether we need to make an y adjustments to our grassroots FCC postreater capital cost estimate for our final rule analy sis. We were not able to make a similar capital cost comparison for the revamp ed FCC postreater capital cost . However, instead of solel y using rule

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- thumb ty pe estimates for revamp costs , we obtained specific cost estimates from the vendors , and thus the capital changes are t echnology

- specific and are dependent on the ty pes of inve stment s made for Tier 2. One advantage of obtaining cost information directly from the vendor companies is that they have information about how the FCC postreaters were configured when they were installed for Tier 2 and how those units are currentl y operated . In addition, they know their technology and would know what it would take to revamp those units for complying with Tier 3. Reviewer s 2 and 3 ’s comment s that vendor 1’s capital cost for revamping an existing FCC postreater were inconsistent with the unit’s severity appear to stem from a lack of understanding of the implications of the Tier 2 baseline and a focus solely on capital costs instead of looking at overall costs

. Based on our request for information, Vendor 1 prov ided cost information for the scenario where the FCC naphtha exiting the FCC unit is 2 0 00 ppm, 800 ppm or 200 ppm and the existing Tier 2 postreaters must desulfurize the sulfur level of FCC naphtha down to 75 ppm. Based on the way

that refiners in each o f these situations invested in Tier 2, the vendor provided one or more way s that refiners in each situation could comply with Tier 3. Vendor 1 reported that in cases when refiners have onl y 200 ppm sulfur in their FCC naphtha they had used minimal capital investments to comply with Tier 2, so they would need to make comparativel y high investments for Tier 3 as they would need to install a new distillation column in addition to add ing catal ys t to the reactor for Tier 3 to avoid large octane losses, high hydrogen consumption and greatl y shortened catalyst life. On the other hand, r efiners with FCC naphtha sulfur levels at 800 ppm invested much more capital for Tier 2 and therefore their situation for compl yi n g with Tier 3 would entail very

modest capital cost investments. We discussed this with Vendor 1 numerous times and adopted it for our anal ys is. The result is that capital costs were a function of Tier 2 investments, and not simply a function of FCC nap htha sulfur levels. Through this process we also discovered that one needs to look at total costs not just capital costs by itself, since when capital costs are low, operating costs tend to be higher , and vice

- versa . Confirming this, however, is very

diff icult because we don’t have access to refinery

- specific capital investment information for Tier 2 ; nor do we have access to actual unit operating information that the vendor companies have access to. We can , however, do what Reviewer s 2 and 3 suggested wh ich is to ask the vendor to review the information provided to us and verify that it correctly represents what the vendor intended. We intend to do that for the final rule analy sis as well as incorporate an y other information available to us that might be tter inform the capital costs. CommentsReviewer s 1 , 2 and 3 suggested requesting additional information from the various

vendors. In some cases, the vendor provided information for one scenario, but not another (such as a grassroots FCC postreater desulfurizing FCC naphtha down to 10 ppm, but not to 25 ppm, thus requiring EPA to estimate the information for the other scenario using other available data). Two vendors did not provide an y information for revamping their existing Tier 2 FCC postreaters to compl y with Tier 3. Response We intend to make additional requests of the vendors so that we may be able to obtain that information and use it for the final rule anal ys is

. C. Cataly st Costs Comments Reviewer s 2 and 3 noticed that Vendor 1 did not provide a catal ys t cost estimate for the grassroots FCC postreater. Reviewer s 2 and 3 also observed that for a revamp scenario that Vendor 1 provided information for, that the vendor had estimated a catal yst cost of 25 to 35 percent of the ISBL capital cost for that unit. Response Although Vendor 1 did not provide a means to estimate catal ys t cost for a grassroots FCC postreater, we in fact estimated its cost and included that cost in our Tier 3 cost estimate. One input fo r estimating cataly st costs was to look back at the catal ys t costs e stimated for Tier 2 and adjusting those costs to current year costs. We also looked at the data provided by other vendors for their catal ys t costs. Since we received the peer review comm ents, we requested Vendor 1 to provide catal ys t cost information for a grassroots unit for Tier 3. Vendor 1 estimated that catal ys t cost would be about 2

– 4 percent of IS BL cost. The cataly st cost that we estimated was 4 percent of ISBL capital cost, so our catal ys t cost is at the upper end of the range provided by Vendor 1. We also asked Vendor 1 about the cataly st cost information provided for that specific revamp case which was provided as a range of 25 to 35 percent ( this was for the addition of a s econd stage for the scenario where FCC naphtha is 2400 ppm exiting the FCC unit) . Vendor 1 stated that those percentages represent the increased catal ys t volume over current catal ys t volume in the Tier 2 reactor, not catal ys t cost estimate as a percent of ISBL. Our catalyst cost estimate is consistent with the way that Vendor 1 intended it. We will further review our catal ys t cost estimate for the final rule cost anal ys is and make adjustments as necessary

. D. Natural Gas Prices CommentsReviewer s 2 and 3 compared the natural gas prices that we used in the refinery

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refinery cost model with their assessment of natural gas prices. The following table contains the comparison of natural gas prices that Reviewer s 2 and 3 provided in their report (the natura l gas prices that we used are found in the column labeled “EPA 2017,” and the those estimated by Reviewer s 2 and 3 are found in the column labeled “Revised 2017”)

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**Natural Gas Prices Comparison ($2009/million Btu)** EPA Base 2009 EIA 2010/2011 EPA 2017 R evised 2017 PADD1 4.87 9.96 7.88 10.30 PADD2 4.86 7.44 7.86 7.87 PADD3 3.49 5.21 5.64 5.64 PADD4 4.36 6.34 7.06 6.77 PADD5 ex CA 4.60 8.33 7.44 8.76 PADD5 CA n.a. 7.04 7.44 7.49

Response Based on the increased natural gas prices in PADDs 1 and 5 outside of California, Reviewer s 2 and 3 estimated that our estimated costs for sulfur control would increase b y 0.02 cents per gallon if we were to use their revised natural gas prices. However, Reviewer s 2 and 3 also suggested that we update our n atural gas prices to those estimated in AEO 2011. Reviewer s 2 and 3 are correct in that we inadvertently based our natural gas prices on the 2010 AEO. The 2011 AEO estimates that natural gas prices in 2017 will be about 25% lower than those in the AEO 20 10 for the ye ar 2017 . However, for the final rule anal ys is we plan on using natural gas prices as estimated in the most recent AEO

. The early release version of the 201 3 AEO projects natural gas prices will be slightly higher in 2017 than those projected in the 201 0 AEO, so if this projection continues in the final version of the 201 3 AEO, the updated natural gas prices will tend to very

slightl y increase our estimated desulfurization costs. E. Hydrogen Demand and Costs Comment Because the h ydrogen demand for some of the desulfurization scenarios provided by the vendor s seemed low, Reviewer s 2 and 3 expressed some concern that the hydrogen consumption values provided b y the vendors reflect stoichiometric values, not actual hydrogen consumption. Reviewer s 2 and 3 stated that if the values reflect stoichiometric hydrogen consumption, then the hydrogen consumption values should be adjusted to reflect actual h ydrogen consumption. The comments by Reviewer s 2 and 3 were made in reference to information provid ed by Vendor 1.Reviewer s 2 and 3 reviewed our h ydrogen cost estimates and concluded that our

estimated hydrogen costs are in line with other estimates , and are appropriate given the natural gas prices that we are using

. Response We contacted Vendor 1 and Vendor 3 since we learned of these comments by Reviewer s 2 and 3 , and b oth vendors confirmed that the h ydrogen consumption values provided to us represent actual h ydrogen consumption including any losses, not stoichiometric hydrogen consump tion only . For the revamp cases in particular, this makes sense. We took a step back to review hydrogen consumption by hydrotreaters. In addition to the hydrogen consumed for desulfurization and the olefin saturation side reactions, there are also losses. One loss of hydrogen during hydrotreating is solution losses. Solution losses are the losses of hydrogen because the hydrogen is dis solved in the FCC naphtha

– the FCC naphtha enters the h ydrotreater with no hydrogen dissolved in the FCC naphtha, but it exits the hydrotreater with hydrogen dissolved in the FCC naphtha. However, for Tier 3, most refiners will be revamping an existing h ydrotreater and the FCC naphtha will already have been treated in the Tier 2 h ydrotreater and therefore already contain hydrogen in solution. Thus, there sh o uld be little or no increased hydrogen loss in solution for the revamped hydrotreaters. Another category

of hydrogen losses during hydrotreating is purge losses. Purge losses are required to prevent the accumulation of hydrogen sulfide gas and other gasses in the recycle hydrogen. However, the very small additional sulfur reduction required for Tie r 3 (reducing FCC naphtha from an average of 75 ppm down to 25 ppm) is not likely to require an y appreciable increase in the purging of recycle gas

. The last category

of h ydrogen losses is mechanical losses. In this case , the hydrogen under pressure leaks out of joints in the pipes and reactors which contain the hydrogen. Because gasoline desulfurization occurs at modest pressures, mechanical losses of hydrogen would be modest in this situation. Furthermore, in the case of revamps, there is little to no additional chance for h ydrogen to leak because in man y cases the Tier 3 revamp relies on the same hydrogen piping and reactor arrangement of the Tier 2 postreater. One reason wh y Reviewer s 2 and 3 suggested that the hydrogen consumption values might represent stoichiometric hydrogen demand , not total hydrogen demand, is because of one of Vendor 1’s hydrogen consumption estimates which seemed low

- onl y 6 standard cubic feet of hydrogen per barrel of feed (SCF/bbl)

. A second hydrogen consumption esti mate, of 19 SCF/bbl was also lower than the other revamp scenarios, which ranged from 30 to 100 SCF/bbl. However, in the context of the revamps which were studied b y Vendor 1, the very low h ydrogen consumption values make sense. In both those cases, the 6 SCF/bbl and 19 SCF/bbl, Vendor 1 indicated that refiners relied on limited FCC postreater investments for Tier 2, choosing to have higher hydrogen consumption to minimize the capital investments needed. However, whenfaced with a more stringent gasoline sulfur standard, refiners would likely choose to make a modest capital investment in lieu of being faced with even larger h ydrogen consumption demands

. In the case of 6 SCF/bbl, the feed to the Tier 2 FCC postreater would only contain a bout 200 ppm sulfur. As discussed above, most r efiners in this situation apparentl y avoid ed the capital cost of installing a splitter to separate out the olefin

- rich light FCC naphtha from the heav y FCC naphtha, choosing instead to feed the entire FCC naphtha stream to the postreater. This resulted in somewhat high er hydrogen consumption for th eir Tier 2 treater. However, instead of avoiding further capital investments for Tier 3 resulting in even higher hydrogen consumption levels (and octane losses and much shorter catal ys t life ) , refiners would likely install a splitter to separate out the l ight FCC naphtha and not run that volume through the hydrotreating reactor. If the refiner were to add the splitter to its Tier 2 FCC postreater but still comply with 30 ppm Tier 2 standard , the refiner would likely reduce the h ydrogen consumption for compl yi ng with Tier 2 (i.e., 20 SCF/bbl ) . Then by increasing the severity

to the Tier 2 treater with splitter to then comply with Tier 3, the FCC postreater might demand an additional 26 SCF/bbl, however, the net h ydrogen increase over their Tier 2 FCC postreater without the splitter would only

be 6 SCF/bbl. A similar story applies for the 19 SCF/bbl revamp case, except that in this case the capital investment that could be made to result in incrementally low hydrogen consumption is a second stage. Additio nal discussion about the investments being made for Tier 3 is provided in Chapter 4 of the draft RIA. F. Electricity

Prices Comment Reviewer s 2 and 3 commented that EPA had used electricit y prices which are low. Reviewer s 2 and 3 stated that EPA’s electricity prices fell in the range of 4.7 to 8.4 c/kw

- hr

. H owever the AEO reports an average electricity price of 6.8 c/kw

- hr for 2009. Response Relative to 2009’s electricity costs for industrial users, our electricity prices may be low

. However, in AEO 2011, EIA estimated that average electricity prices in 2017 for industrial users will average 6.0 c/kw

- hr , which is in the middle of our range. For our final rule anal ys is, we will use the EIA’s most recent AEO for estimating electricity prices to refiners

. The early release of the 201 3 AEO shows a higher average electricity price for 2017 of 7.0 cents per kilowatt

- hour, and if the latest AEO maintains this electricity price, it will slightly increas e our desulfurization cost estimates. As Reviewer s 2 and 3 acknowledged though , small differences in electricity prices would have a negligible impact on the estimated desulfurization costs. G. Octane Costs CommentReviewer s 2 and 3 provided some context for octane costs

. Reviewer s 2 and 3 note that

the regular grade

- premium grade price differential between 2008 and 2010 averaged $5.29/gal, or $0.90/bbl per octane number. Furthermore, as Reviewer s 2 and 3 point out, the historical methods for estimating the cost of making up lost oct ane , such as the regular

- premium price differential, do not account for increased blending of ethanol into the future. Reviewer 2 and 3 stated that its future year octane costs based on its World model are similar to ours. They estimated costs for octane which ranged from 0.25 to 0.45 $/octane

- bbl for PADD 3 and PADD 1, respect ivel y (which is 0.59 to 1.07 cents/octane

- gallon). Response We agree with the comments by Reviewers 2 and 3 about the disadvantage of using historical regular grade

- premium grade price differential method for estimating octane cost because it doesn’t consider future use of ethanol, and also we believe is that it likely includes some of the profit earned by the much higher premium gasoline prices . Also, part of the price premium f

or premium gasoline could be due to the need to segregate and distribute a much smaller volume of product. In our case, we are solely interested in the cost of making up losses in gasoline octane within the refinery , not the octane cost difference of high octane blends at retail

. For this case, LP refinery models are better suited for estimating octane costs because they solel y reflect the cost for recovering the lost octane. Reviewer 2 and 3 stated that future ye ar octane costs based on its World mode l are similar to ours. They estimated costs for octane which ranged from 0.25 to 0.45 $/ octane
- bbl for PADD 3 and PADD 1, respect ively (which is 0.59 to 1.07 cents/octane

- gallon). Our estimated octane cost developed using the Haverl y LP refinery cost mod el is 0.32 $/octane

- bbl (0.76 cents/octane

- gallon) is right in the range of costs estimated by Reviewer s 2 and 3

. Reviewer 2 and 3 ’s refinery

model runs were run in 2015 while our refinery modeling runs were made in 2017. Because Reviewer s 2 and 3 refer to E10 runs, their runs were likely conducted with about 100% E10 in the gasoline pool. Our runs assuming 2017 renewable fuels volumes resulted in about half E10 and half E15 in the gasoline pool. Reviewer s 2 and 3 also found that the more ethanol that is being blended into the gasoline pool, the lower the marginal cost for making up for lost octane , and our analy sis agrees with that statement. We chose, however, to assess the cost of recovering lost octane on the first ye ar of the proposed Tier 3 gasol ine desulfurization program, which is 2017, instead of later ye ars when projected octane costs would likely be lower , probabl y making our octane cost projections , based on the first ye ar of the Tier 3 program, somewhat conservative. Reviewer s 2 and 3 comm ented that our octane costs were not lowest in PADD 3 (Gulf Coast), which ty pically is the lowest refining cost PADD of refineries. Because of this concern about the relative projected octane recovery costs in the different PADDs, w e will review and possibly rerun our octane cost cases for the final rulemaking (FRM)

. H. Li ght Straight Run (L SR) TreatingComment

Reviewer s 2 and 3 commented that the refinery

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- refinery cost model did not appear to apply an overdesign factor to L SR treaters. Response Initiall y a dialog box was created that contained a 1.15 overdesign factor to apply for LSR treating. However, the overdesign factor was subsequently integrated directly into the cost equation for LSR treating , hence the 1.15 factor contained in the dialo g box was never utilized , although an overdesign factor was included

. I. Other Cost Inputs Comment Reviewer s 2 and 3 commented that the cost model did not appear to consider cost factors outside of FCC naphtha and L SR desulfurization units, such as acid gas and sulfur recovery plants. Response The cost for operating s ulfur recovery plants w as estimated and the very low cost impact wa s included in the refinery cost model. J. Future Refining Industry Changes Impacting Desulfurization Costs Comments Reviewer s 2 and 3 commented that over time changes in the US refining i ndustry should be accounted for because of their likely impact on gasoline desulfurization costs. These include

 Accounting for major projects at US refineries particularl y those which will allow refineries to process volumes of heav y Canadian crudes (tar sands) , including the potential impact of the Key stone pipeline if it were built.

 Potential shifts over time in FCC feedstocks and operating modes, notabl y: shift towards more resid

ual fuel in FCC feed as VGO is increasingl y “pulled away” to hydro

- cracking units as demand for distillates grows, FCC cataly st/operating shifts to lower gasoline higher distillate (LCO) and also higher propane yi elds; potential resulting impacts on FCC gasoline q

ualities, volumes and thus concentration in the U.S. gasoline pool

 Potentially

substantial changes in U.S. domestic oil liquids supply , both light sweet crudes from the Bakken , Eagle Ford, Niobrara, Utica shale and other emerging domestic crude oil supplie s and the potential impacts of rising NGL’s supply , and Extent to which current trends for growth in U.S. refined products exports will continue

and their potential impacts on U.S. refining. Response Reviewer s 2 and 3 raise d some important points about ho w future US refining industry changes could impact the Tier 3 modeling effort. For the NPRM cost an a ly sis, we initially identified that three US refineries were to undergo major crude oil throughput expansions. To obtain information about the expansions, we reviewed the unit capacit y information from EIA for 2011 (our base ye ar was 2009 so we originally relied on 2009 refinery

capacity data from EIA and the Oil and Gas Journal) and used the later data which included the in creased crude and other unit thro ughput increases due to the expansions

. Reviewer s 2 and 3 provided information about the refinery

throughput expansion or p rojects that would allow refineries to process Canadian tar sands or other heav y crude oil. W e intend to in corporate the refinery expansion/modification information provided b y Reviewer s 2 and 3 and any we can find from other sources in our final rule cost anal ys is to the extent possible within the constraints of a refinery

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model

. K. Changes in Crude Quality Comment Revie wer s 2 and 3 also pointed out that crude quality will change over time which would impact gasoline desulfurization costs

. Response Very sour (high sulfur) Canadian bitumen or tar sands, as well as sweet (low sulfur) crude oils from shale are increasingl y being processed by US refineries. Because of the very different nature of Canadian tar sands, refineries planning on processing it typically undergo significant capital modifications and we therefore know fairl y confidently which re fineries will be processing tar sands in the future, which is helpful when running a refinery

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- refinery cost model. However, predicting how refineries will process the tar sands provides an additional challenge. We will attempt to figure out how to int egrate the increased use of tar sands in our cost anal ys is. The development of shale resources can yi eld both shale oil and light hydrocarbons such as natural gas liquids. Shale oil is ty pically refined, and natural gas liquids are t ypicall y blended b y the refineries which are located nearb y the location where these hydrocarbon sources are being developed. We will assess how to include these expanding sources of h ydrocarbons in our cost anal ys is for the final rule within the constraints of a refinery

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model

. L. In creased Dieselization CommentsReviewer s 2 and 3 su ggested that we account for the change in refinery operations caused by the shifting product market which increasingl y favors diesel fuel over gasoline

. Revieweres 2 and 3 pointed out that to respond to these changing product demands, refiners could shift vacuum gas oil, the t ypical feedstock for FCC units , to hydrocrackers resulting in heavier, more sour feedstocks for FCC units which could increase desulfurization costs. Response Since Tier 3 will take effect in the future, our cost anal ys is should take into account, as much as possible, future trends in refining. As diesel demand increases while gasoline demand lags, refiners are expected t o add hydrocracking capacit y which favors diesel fuel production

. While refiner s may

make the feedstock changes to FCC units that Reviewers 2 and 3 suggest, this change to the FCC would be offset b y the increased production of low sulfur hydrocrackate which would displace so me of the volume of FCC naphtha blended into the gasoline pool, thus tending to lower the cost of desulfurization. We will consider way s to incorporate these changes for the final rule. H owever, it is challenging to make such adjustments in a refinery

- sp ecific model since it is difficult to predict the refineries that will add hydrocracker capacity

. Another likely and important change of dieselization is the undercutting of FCC naphtha into the distillate pool. I f refiners undercut their FCC naphtha into distillate, it would reduce the volume of FCC naphtha needing to be h ydrotreated thus decrease the volumetric demand on the existing Tier 2 postreater. However, more importantly , undercutting FCC naphtha can reduce by half the sulfur exiting the FCC unit. Overall, these changes would reduce the sulfur and volume content of the feed to the existing postreaters signific antly easing compliance with Tier 3 even if refiners reserve the dirtiest of feeds to the FCC unit

. M. Refinery

- by

- Refinery Model Documentation and Organization Comments Reviewers 1 , 2 and 3 commented that the refinery

- by

- refinery cost model lacked doc umentation , needed simplified formula structure and was poorly laid out which made navigating the refinery

cost model difficult. Despite the difficulty navigating the refinery

cost model, both reviewers indicated that they had adequate time to review the refinery

- by

- refinery cost model. Response Our plan is to heed the Reviewers’ advice b y organizing and better document ing the r efinery

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refinery cost model

.