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OPPORTUNITIES FOR 

BIORENEWABLES IN OIL REFINERIES

FINAL TECHNICAL REPORT

SUBMITTED TO: 

U.S. DEPARTMENT OF ENERGY

uop

25 East Algonquin Road 

Des Plaines, Illinois 60017-5017 

U.S.A. 

DOE Award Number: DE-FG36-05GO15085 

Period of Performance: 12/01/2004 to 04/15/2005 

Technical Point of Contact: Richard Marinangeli 

Manager, New Directions Exploratory and Fundamental Research 

UOP LLC 

(847) 391 3327 (phone) 

(847) 391 3724 (fax) 

richard.marinangeli@uop.com 

Contributors 

UOP Terry Marker, John Petri, Tom Kalnes, Mike McCall, Dave Mackowiak, Bob 

Jerosky, Bill Reagan, Lazlo Nemeth, Mark Krawczyk 

NREL Stefan Czernik 

PNNL Doug Elliott 

Michigan Technological University David Shonnard 
Opportunities for Biorenewables in Oil Refineries 

 

TABLE OF CONTENTS

I Executive Summary................................................................................................ 1

II Study Goals and Basis........................................................................................... 3

II.1 Study Goals.....................................................................................................................3

II.2 Study Basis .....................................................................................................................3

II.2.1 Market and Supplies ...................................................................................................3

II.2.2 Feedstock costs and product prices ...........................................................................4

II.2.3 Refinery Basis.............................................................................................................5

III Refinery Opportunities for Vegetable Oils and Grease....................................... 8

III.1 Biodiesel (Fatty Acid Methyl Ester) Production...............................................................8

III.1.1 Fatty Acid Methyl Ester Production in Existing Refining Units ...............................9

III.1.2 Retrofit Enablers...................................................................................................11

III.1.3 Conclusions ..........................................................................................................13

III.2 Green Diesel Production...............................................................................................14

III.2.1 Green Diesel Production Approaches..................................................................14

III.2.2 Green Diesel Production – Co-Processing Biofeedstocks in an Existing Unit.....14

III.2.3 Green Diesel Production – Processing Biofeedstocks in a Stand-Alone Unit......18

III.2.4 Green Diesel Experimental Results .....................................................................20

III.3 Comparison of Biodiesel and Green Diesel Process and Properties ...........................21

III.4 Green Gasoline and Green Olefins Production ............................................................22

III.5 Economic Assessment and Life Cycle Analysis ...........................................................23

IV Refinery Opportunities for Pyrolysis Oil ............................................................ 25

IV.1 Background and Processing Options............................................................................25

IV.2 Gasoline And Aromatics Production from Pyrolytic Lignin............................................27

IV.3 Coprocessing with VGO in an FCC...............................................................................29

IV.4 H2 production from Pyrolysis Oil...................................................................................31

IV.4.1 Integration of Pyrolysis Oil Reforming into Refinery Hydrogen Plants.................31

IV.4.2 Aqueous Phase Reforming (APR) of the Aqueous Phase of Pyrolysis Oil..........33

IV.5 Hydrothermal Processing and Blending Studies...........................................................35

IV.6 Economic Analysis of Pyrolysis Oil Processing Options...............................................36

V Opportunities for Fischer Tropsch Liquids........................................................ 39

VI Acronyms and Abbreviations.............................................................................. 41

VII References ........................................................................................................... 43

i
Opportunities for Biorenewables in Oil Refineries 

 

LIST OF FIGURES

Figure 1. Global feedstock supplies. (a) Availability of oil and grease relative to other 

fuels.,

,

,

 (b) Potential contribution to chemicals production.,...................................... 4

Figure 2- Average feedstock prices. Note that $306/ton = $40/bbl for WTI crude (2nd 

quarter 2004)............................................................................................................. 4

Figure 3. Biorenewable and petroleum feedstock prices. ................................................ 5

Figure 4. Typical U.S. refinery........................................................................................... 6

Figure 5. Typical U.S. hydrocracking refinery. .................................................................. 7

Figure 6. Vegetable oil and grease processing routes...................................................... 8

Figure 7. UOP Merox processing schemes. ..................................................................... 9

Figure 8. Possible Merox unit retrofit to FAME service................................................... 10

Figure 9. UOP MTBE (Ethermax) technology................................................................. 11

Figure 10. Alternatives for green diesel production......................................................... 14

Figure 11. Effect of co-processing brown grease on catalyst temperature cycle............ 16

Figure 12. Effect of co-processing brown grease on catalyst deactivation rate. ............. 16

Figure 13. Effect of co-processing brown grease on catalyst cycle length. .................... 17

Figure 14. Proposed flow scheme for standalone production of green diesel. ............... 18

Figure 15. Decarboxylation vs Hydrodeoxgenation ........................................................ 19

Figure 16. FCC processing approach. ............................................................................ 23

Figure 17. Capital costs of biofuels production from oils and greases............................ 23

Figure 18. NPV comparison of biofuels and chemicals................................................... 23

Figure 19. LCA – single environmental impact score ..................................................... 24

Figure 20. LCA –comparison of climate-active CO2....................................................... 24

Figure 21. Typical pyrolysis process............................................................................... 25

Figure 22. Structure of lignin........................................................................................... 26

Figure 23. Pyrolysis oil processing routes....................................................................... 27

Figure 24. Process flow for hydrotreating/hydrocracking of pyrolytic lignin. ................... 29

Figure 25. Adiabatic pre-reformer for the aqueous-phase of pyrolysis oil....................... 32

Figure 26. Proposed APR process flow scheme............................................................. 34

Figure 27. Years to payback for conversion of pyrolytic lignin to gasoline...................... 37

Figure 28. Years to payback for combined process........................................................ 38

Figure 29. Typical hydrocracking refinery processing crude FT liquids. ......................... 40

ii
Opportunities for Biorenewables in Oil Refineries 

 

LIST OF TABLES

Table 1. Availability of biorenewable feedstocks in U.S.,,,, ................................................ 3

Table 2. Summary of FAME experimental results........................................................... 12

Table 3. Hydrotreating yields for brown grease. ............................................................. 15

Table 4. Hydrotreating free fatty acids. ........................................................................... 15

Table 5. Hydrodeoxygenation versus decarboxylation theoretical yields........................ 19

Table 6. Green diesel experimental summary. ............................................................... 20

Table 7. Biodiesel and green diesel properties ............................................................... 21

Table 8. Comparison of biodiesel and green diesel production processes..................... 21

Table 9. Green diesel and biodiesel yields. .................................................................... 22

Table 10. Green olefin yields. ......................................................................................... 22

Table 11. Green gasoline yields...................................................................................... 22

Table 12. Material balance for pyrolysis oil production. .................................................. 26

Table 13. Pyrolysis oil properties. ................................................................................... 26

Table 14. Results from hydrotreating pyrolytic lignin. ..................................................... 28

Table 15. Hydrotreating+hydrocracking pyrolytic lignin. ................................................. 28

Table 16. Yield estimate for hydrotreating/hydrocracking pyrolytic lignin. ..................... 28

Table 17. Yields for FCC conversion of pyrolysis oil, pyrolytic lignin and HT pyrolysis oil.29

Table 18. Carbon residue results.................................................................................... 30

Table 19. Product properties of total liquid products from FCC testing at 1000F and 5/1 

cat/oil ratio............................................................................................................... 30

Table 20. Composition of aqueous-phase compounds from pyrolysis oil ....................... 31

Table 21. Pre-reformer syngas composition for the aqueous phase of pyrolysis oil....... 31

Table 22. Reforming equilibrium reactions for APR of pyrolysis oil components........... 33

Table 23. Comparison of APR process flow schemes................................................... 34

Table 24. Comparison of H2 production routes. .............................................................. 35

Table 25. Hydrothermal processing experiment conditions. ........................................... 35

Table 26. Product properties of hydrothermally processed pyrolytic lignin..................... 36

Table 27. Fuel value of pyrolysis oil................................................................................ 36

Table 28. Incentives for gasoline production from pyrolytic lignin.................................. 37

Table 29. Incentives for H2 production ............................................................................ 38

Table 30. Overall incentives for use of pyrolysis oil for fuel and gasoline production.... 38

Table 31 Properties of crude FT liquids .......................................................................... 39

Table 32. Detailed properties of crude FT liquids. .......................................................... 39

Table 33. Value of crude FT liquids to refiners. .............................................................. 40

Table 34. Acronyms and abbreviations........................................................................... 41

iii
 

I EXECUTIVE SUMMARY

The purpose of this study was to evaluate the potential for using biorenewable feedstocks in oil 

refineries. Economic analyses were conducted, with support from process modeling and proof of 

principle experiments, to assess a variety of potential processes and configurations. The study 

identified a number of promising opportunities for biorenewables in existing or new refining 

operations. 

The work was performed as a joint project with the National Renewable Energy Laboratory 

(NREL), and Pacific Northwest National Laboratory (PNNL). NREL provided pyrolysis oil 

feedstocks and biorenewable expertise. PNNL provided feedstocks and biorenewable processing 

expertise; PNNL also conducted some of the proof of principle experiments. The DOE provided 

$80,000 of the total $230,000 funding for the project. 

The study considered two primary alternatives: the production of biodiesel and green diesel from 

vegetable oils and greases and opportunities for utilization of pyrolysis oil. Green diesel produced 

from hydrotreating vegetable oil and grease was identified as one of the best options for refiners, 

agricultural producers, and forest products producers. Green diesel has superior product 

properties, requires less capital investment to produce, and has a lower environmental impact than 

biodiesel. Green diesel is economically attractive under two conditions: (1) using a low-cost 

feedstock such as brown grease; or (2) if vegetable oil feedstocks are subsidized.. Green gasoline 

and green olefins, which are produced by cracking vegetable oils and grease in an FCC unit, are 

also economically attractive, particularly if eligible for subsidies. 

Processing of pyrolysis oils represents a longer term opportunity, since pyrolysis oils are 

currently available only in small quantities. However, unlike vegetable oils and greases, they have 

the potential to provide a significant portion of future transportation fuels (60%). The study 

evaluated options for using crude pyrolysis oil as well as two of its fractions – pyrolytic lignin 

and the water-soluble phase . 

Experimental work confirmed that pyrolytic lignin can be converted to gasoline with reasonable 

H2 consumption using mild hydrotreating followed by hydrocracking. The economics of this 

process are increasingly attractive with increased crude oil price or decreased pyrolysis oil price 

with increased production. At a price of $50/bbl for crude, the analysis predicted payback in three 

years. 

The water-soluble phase of pyrolysis oil is suitable for generating hydrogen. Most, or all, of the 

hydrogen used in a typical refinery can be provided by steam reforming 5200 bpd of the water 

soluble portion of pyrolysis oil. A small portion of this hydrogen can be used to provide hydrogen 

for hydrotreating the pyrolytic lignin to make gasoline. The economics of hydrogen formation 

from the water soluble portion of pyrolysis oil become attractive when there is a CO2 credit of 

$30/ton and pyrolysis oil cost is $10/bbl. Since the pyrolysis oil used for hydrogen production 

would replace purchased natural gas, this is a way to reduce nearly 20% of refinery emissions of 

climate-active CO2. However, pyrolysis oil price must drop before this process becomes 

economically attractive. 

Crude pyrolysis oil can be substituted for purchased natural gas as a fuel in refinery boilers and 

furnaces, also reducing CO2 emissions. Pyrolysis oil is less costly than natural gas by 

$1.2/MMBTU and therefore represents an economically viable substitute for natural gas in a 

refinery setting. 

The high oxygen content of pyrolysis oil imposes constraints on processing in standard 

metallurgy and current economics for hydrotreating pyrolysis oil assume a standalone unit is 

required. Initial hydrothermal treatment under mild conditions may provide a low cost approach 

1
 

to reduce the oxygen content and allow processed in standard equipment. More experimental 

work is needed to determine the viability and economic impact of hydrothermal treatment. 

Hydrotreating and hydrocracking, technologies which were developed for oil refineries, were 

found to be particularly useful for processing vegetable oil, greases, and pyrolysis oils. The level 

of hydrogen required for hydrotreating can be controlled by processing conditions and catalyst, 

since the same materials which catalyze hydrodeoxygenation also catalyze decarboxylation. 

Optimization of this technology for biorefineries will lead to better utilization of these 

biofeedstocks for conversion to fuels and chemicals. 

A key barrier to the use of biorenewables in oil refineries is that available quantities are small 

relative to the size of most refinery processing equipment. Refiners are naturally resistant to 

risking catalysts and run time in large petroleum processing units, for the sake of processing small 

amounts of biofeedstocks. Skid mounted, standalone green diesel, or pyrolysis oil hydrotreating 

units solve this problem by allowing biorenewables to be processed separately from petroleum 

fractions. These skid mounted units could also be used by agricultural and forest producers. 

2
 

II STUDY GOALS AND BASIS

II.1 STUDY GOALS

The overall goal of this study was to evaluate the economics of using biomass-derived feedstocks 

(biofeedstocks) in petroleum refineries. The technical work was divided into three tasks: 

modeling of refinery options to project yields and costs (Task 1); a variety of scoping 

experiments (Task 2); and, analysis of costs and process economics (Task 3). 

Several alternatives were evaluated: 

• Processing of biofeedstocks in existing refinery equipment 

• Coprocessing of biofeedstocks with crude oil fractions in existing refinery equipment 

• Blending biofeedstocks with petroleum feeds 

In addition, new processing steps were evaluated that would enable more effective use of 

biofeedstocks in refining operations. 

Two classes of biofeedstocks were considered: 

• Vegetable oils and greases derived from crops such as corn, soy, canola and rapeseed 

• Pyrolysis oil derived from forest waste 

A number of processing options was considered for each type of biofeedstock. In each case, one 

or more candidate processing schemes was modeled. Experiments were conducted as needed to 

determine the key processing parameters required to understand cost and economic impact of 

biofeedstock use in a particular process. Cost estimates were developed and used in a broader 

evaluation of the potential process economics. This report summarizes the results of the process 

models and supporting experiments. Additional details are provided in the supporting Appendix. 

This Final Report is organized by processing option, rather than by task, to more clearly present 

the alternatives. The basis and assumptions used in the study are outlined in Section II.2. Section 

III describes opportunities for utilizing vegetable oils and greases, particularly in the production 

of biofuels. Section IV describes the evaluation of opportunities for processing or coprocessing 

pyrolysis oil to produce gasoline, aromatics or hydrogen. The value of processing FischerTropsch liquids in biorefineries was also evaluated, described in Section V. Lists of acronyms and 

abbreviations and key references are located in Sections VI and VII, respectively. The Appendix 

provides a summary of key experiments (Appendix Tables 1 and 2) and additional background 

for the study basis. 

II.2 STUDY BASIS 

II.2.1 Market and Supplies 

Table 1. Availability of biorenewable feedstocks in U.S.1, , , , 2 3 4 5

3

Definition 

U.S. Production 

(bpd) 

Available for 

Conversion to 

Fuels (bpd) 

Vegetable oils Produced from soybeans, corn, 

canola, palm 

194,000 33,500 

Recycled 

products 

Yellow grease, brown (trap) grease 51,700 33,800 

Animal fats Tallow, lard, fish oil 71,000 32,500 

Pyrolysis oil Produced by pyrolysis of wood 

slash, waste and other biomass 

waste 

1,500 750 
 

50000

0

10000

20000

30000

40000

Current Current

Liquid Transport Fuels Liquid Transport Fuels

Diesel

Available Oil/Grease Available Oil/Grease

Potential Potential

Cellulosic Cellulosic Waste

kbpd

100

0

10

20

30

40

50

60

70

80

90

Ethylene Ethylene Propylene Propylene Available Oil Available Oil

and Grease and Grease

M ton/yr M ton/yr

(a) (b)

50000

0

10000

20000

30000

40000

Current Current

Liquid Transport Fuels Liquid Transport Fuels

Diesel

Available Oil/Grease Available Oil/Grease

Potential Potential

Cellulosic Cellulosic Waste

kbpd

50000

0

10000

20000

30000

40000

Current Current

Liquid Transport Fuels Liquid Transport Fuels

Diesel

Available Oil/Grease Available Oil/Grease

Potential Potential

Cellulosic Cellulosic Waste

kbpd

0

10000

20000

30000

40000

Current Current

Liquid Transport Fuels Liquid Transport Fuels

Diesel

Available Oil/Grease Available Oil/Grease

Potential Potential

Cellulosic Cellulosic Waste

Potential Potential

Cellulosic Cellulosic Waste

kbpd

100

0

10

20

30

40

50

60

70

80

90

Ethylene Ethylene Propylene Propylene Available Oil Available Oil

and Grease and Grease

M ton/yr M ton/yr

100

0

10

20

30

40

50

60

70

80

90

Ethylene Ethylene Propylene Propylene Available Oil Available Oil

and Grease and Grease

M ton/yr M ton/yr

0

10

20

30

40

50

60

70

80

90

Ethylene Ethylene Propylene Propylene Available Oil Available Oil

and Grease and Grease

M ton/yr M ton/yr

(a) (b)

Figure 1. Global feedstock supplies. 

(a) Availability of oil and grease relative to other fuels.6,7

,

8,9

 (b) Potential contribution to chemicals production.10,11

Roughly 100,000 bpd of oil and grease are currently available in the U.S. for conversion to fuels 

(Table 1).Oils and grease are available today and are likely to make an increasing contribution to 

the world’s fuel supply in the near term, in part due to government-mandated targets (see 

Appendix, Table 3). For example, the U.S. has mandated that 2.5% of diesel be generated from 

biorenewable sources by 2010. However, quantities of oil and grease are limited. In contrast, very 

little pyrolysis oil is produced at present but the potential supply is much larger. As shown in 

Figure 1 (a), significant levels of cellulosic waste are available globally for production of 

pyrolysis oil if conversion to transportation fuels becomes economically attractive. Therefore, 

pyrolysis oils have a far greater long term potential for replacing transportation fuels. Chemicals 

are a much smaller market than transportation fuels and oils and grease feedstocks could play a 

significant role, particularly for olefin production, as shown in Figure 1.

II.2.2 Feedstock costs and product prices 

The study took into account both feedstock costs and the projected prices of potential products. 

The economic analysis was primarily based on a West Texas Intermediate (WTI) crude feedstock 

price of $40 per barrel. The impacts of price increases to $50 and $60 per barrel of crude were 

also explored. Biofeedstock prices were based upon an average of the 2003 – 2004 values 

reported in the Chemical Marketing reporter (See Appendix, Table 4). Figure 2 shows the trend in 

feedstock prices in $/ton during the 2003 – 2004 period. 

0

100

200

300

400

500

600

700

800

2003

1Q

2003

2Q

2003

3Q

2003

4Q

2004

1Q

2004

2Q

2004

3Q

2004

4Q

WTI crude

Reg gasoline

Diesel

Yellow grease

Soybean oil

Ethylene

Propylene

Methanol

Residual fuel

$/ton

0

100

200

300

400

500

600

700

800

2003

1Q

2003

2Q

2003

3Q

2003

4Q

2004

1Q

2004

2Q

2004

3Q

2004

4Q

WTI crude

Reg gasoline

Diesel

Yellow grease

Soybean oil

Ethylene

Propylene

Methanol

Residual fuel

$/ton

Figure 2- Average feedstock prices. 

Note that $306/ton = $40/bbl for WTI crude (2nd quarter 2004) 

4
 

A number of factors impact the estimated price of biofuels. The following factors were included 

in the analysis: 

• Capital costs. Capital costs were estimated for equipment modifications or for new 

plants, depending upon process requirements. 

• Transportation costs. Transportation costs vary depending on the location of the 

refinery and the location of the biofeedstock. A number of U.S. refineries are located near 

sources of biomass feedstock, specifically oil and grease production facilities and 

agricultural and wood resources (Appendix, Figures 1 and 2). A transportation charge of 

$2 per barrel was assumed for biofeedstocks based on typical rail transportation costs for 

grease and oils. 

• CO2 credits. The use of pyrolysis oil as a fuel would generate a CO2 credit which could 

be traded. Currently CO2 is trading for $10/ton but is expected to reach $30/ton in 5 

years. 

• Subsidies. Potential subsidies were calculated based on the subsidies for biodiesel shown 

in Table 5 of the Appendix. 

• Cetane and octane numbers. Cetane number (see Section III.3) affects the price of both 

green diesel and biodiesel. The value of $0.13/cetane-bbl was used to reflect the impact 

of cetane number on prices; the impact of octane number was represented as 

$0.50/octane-bbl. 

Figure 3 compares the prices of biofeedstocks projected in this study to those of petroleum-based 

feedstocks. While some biofeedstocks are significantly more expensive than petroleum-based 

alternatives, others are expected to be less expensive, particularly if supported with government 

subsidies. 

The assumptions used to project the cost of 

pyrolysis oil are summarized in Table 6 and 

Table 7 of the Appendix. In the best case, the 

projected cost of $4.8/MMBTU for pyrolysis 

oil is significantly higher than that of fuel oil 

($4.2/MMBTU) on a heating value basis. 

Subtracting a $30/ton petroleum feedstock 

price credit would reduce the pyrolysis oil 

cost to $4.1/MMBTU, but inclusion of 

transportation charges increases the cost of 

pyrolysis oil to $4.7/MMBTU, still 

significantly higher than heavy fuel oil. A 

major factor is a capital charge of $81MM for 

a 6800 bpd pyrolysis unit, which boosts 

pyrolysis oil from $9/bbl to $16/bbl in the 

case of a $30/ton feedstock. 

0

10

20

30

40

50

60

70

80

Petroleum Based Bio Based

$/bbl

Residual Oil

WTI/Brent Crude

Soy Oil

Soy Oil w/subsidy

Yellow Grease

Yellow Grease w/subsidy

Brown Grease

Pyrolysis Oil

Petroleum Petroleum 

Based

Bio Based Bio Based

0

10

20

30

40

50

60

70

80

Petroleum Based Bio Based

$/bbl

Residual Oil

WTI/Brent Crude

Soy Oil

Soy Oil w/subsidy

Yellow Grease

Yellow Grease w/subsidy

Brown Grease

Pyrolysis Oil

Petroleum Petroleum 

Based

Bio Based Bio Based

Figure 3. Biorenewable and petroleum 

feedstock prices. 

II.2.3 Refinery Basis 

The study was based upon a standard 150,000 bpd U.S. oil refinery. Two typical configurations 

and process units are pictured in Figure 4 and Figure 5. Almost all refineries contain fluid 

catalytic crackers (FCC’s), reformers and a variety of hydrotreaters. Many refineries contain 

hydrocrackers which crack vacuum gas oil (VGO) to make gasoline or diesel. 

5
 

CH4 H2 Plant H2

C

D

U

Crude

Oil

SGP

V

D

U

FCCU

Coker

Coke

Diesel 

Hydrotreater

NHT

Hydrotreater

in Kerosene Jet

12,000 B/D

150,000 B/D 36,000 B/D 28,000 B/D

Gasoline

ALKY

12,000 B/D

MTBE

Diesel

LPG

Butanes

Gasoline 

Hydrotreater

31,000 B/D

17,000 B/D

26,000 B/D

CSO

F

Fuel Oil

iC4

Reformer

CH4 H2 Plant H2

C

D

U

Crude

Oil

SGP

V

D

U

FCCU

Coker

Coke

Diesel 

Hydrotreater

NHT

Hydrotreater

in Kerosene Jet

12,000 B/D

150,000 B/D 36,000 B/D 28,000 B/D

Gasoline

ALKY

12,000 B/D

MTBE

Diesel

LPG

Butanes

Gasoline 

Hydrotreater

31,000 B/D

17,000 B/D

26,000 B/D

CSO

F

Fuel Oil

iC4

Reformer

Figure 4. Typical U.S. refinery. 

The ability to process biofeedstocks in existing refineries is affected by their high oxygen and 

acid content. For example, oils and greases typically contain 11-12% oxygen while pyrolysis oils 

may contain 35-45% oxygen. The acidity of feeds processed in a crude oil refinery is typically 

measured by the Total Acid Neutralization (TAN) number, defined as the milligrams of 

potassium hydroxide required to neutralize one gram stock. The TAN numbers for grease and 

pyrolysis oil make it difficult to process these materials in standard refinery equipment. However 

they may be processed in 317 stainless steel clad vessels. Other properties of typical 

biofeedstocks are summarized in Table 8 of the Appendix. 

Some of the biofeedstock processing options considered in the study require added hydrogen. 

Almost all refineries include hydrogen plants to supply hydrogen for hydrotreatment to remove 

sulfur from diesel and gasoline. The economic analyses presented here assumed an average U.S. 

hydrogen plant supplying 25 million cubic feet per day of hydrogen using a natural gas feed. 

6
 

HCK

SGP SGP

FCCU

VISB

ISOM

LPG

Butanes

Gasoline

Jet

CSO

Diesel

CH4

Fuel Oil

LCO

NHT

MEROX

C

D

U

Crude

Oil

All H2

Users

MTBE

ALKY

F

F

F

H2 Plant

V

D

U

Diesel 

Hydrotreating

Reformer

HCK

SGP SGP

FCCU

VISB

ISOM

LPG

Butanes

Gasoline

Jet

CSO

Diesel

CH4

Fuel Oil

LCO

NHT

MEROX

C

D

U

Crude

Oil

All H2

Users

MTBE

ALKY

F

F

F

H2 Plant

V

D

U

Diesel 

Hydrotreating

Reformer

Figure 5. Typical U.S. hydrocracking refinery. 

7
 

III REFINERY OPPORTUNITIES FOR VEGETABLE OILS AND GREASE

Five separate processing routes for vegetable oil and grease were compared, including biodiesel, 

green diesel, green gasoline, green olefins and green detergent. The alternatives are illustrated in 

Figure 6. Biodiesel is differentiated from green diesel in the study. Biodiesel is typically 

composed of fatty acid methyl esters (FAME). “Green diesel” refers to an acceptable diesel pool 

blend component produced from a suitable biomass. Green diesel here refers to hydrocarbons 

produced from biofeedstocks that boil in a typical distillate fuel boiling range and do not contain 

oxygen in their molecular structure. 

Biodiesel production is discussed in Section III.1, together with enabling technologies to facilitate 

the use of existing refining facilities. Green diesel production by either coprocessing in existing 

units or processing in new standalone units is evaluated in Section III.2. Biodiesel and green 

diesel are compared in Section III.3. The production of green gasoline and olefins is reviewed in 

Section III.4. This section concludes with an overall assessment of process economics and life 

cycle analysis in Section III.5. 

Biodiesel

Diesel 

Hydrotreater

Gasoline

Methanol

Veg oil/Grease

Biodiesel

Glycerol

Veg oil/Grease

Diesel

Diesel 

+ Green diesel

H2

FCC

VGO

Petro-FCC

VGO

C2=+C3=

Hydrotreater Detergent 

Paraffins

H2

Olefins

Chemicals

Veg oil/Grease

Veg oil/Grease

Green Diesel 

(coprocess or stand alone)

Green Gasoline

Biodiesel

Application Feed Process Product

Veg oil/Grease

Biodiesel

Diesel 

Hydrotreater

Gasoline

Methanol

Veg oil/Grease

Biodiesel

Glycerol

Veg oil/Grease

Diesel

Diesel 

+ Green diesel

H2

FCC

VGO

Petro-FCC

VGO

C2=+C3=

Hydrotreater Detergent 

Paraffins

H2

Olefins

Chemicals

Veg oil/Grease

Veg oil/Grease

Green Diesel 

(coprocess or stand alone)

Green Gasoline

Biodiesel

Application Feed Process Product

Veg oil/Grease

Figure 6. Vegetable oil and grease processing routes. 

III.1 BIODIESEL (FATTY ACID METHYL ESTER) PRODUCTION

Due to environmental regulations and changing requirements in fuel specifications, the 

processing requirements for petroleum products are changing. For example, limits placed on the 

amount of sulfur in motor fuels have led certain refiners to install new hydrotreating capacity 

while in some cases shutting down existing treating units. Such events can provide the refiner an 

opportunity to retrofit existing process equipment, such as storage tanks, pumps, heat exchangers, 

reactors, separators, and columns, to bio-fuel production. The economic viability of such retrofits 

is highly dependent on availability and cost of renewable feedstock(s) and the markets for the 

refined products. The study investigated options and enablers for retrofitting existing units to 

produce fatty acid methyl esters (FAME). 

8
 

III.1.1 Fatty Acid Methyl Ester Production in Existing Refining Units 

The study considered two specific examples of retrofitting existing refining equipment to Fatty 

Acid Methyl Ester (FAME) production from biofeedstocks: 

• retrofitting a UOP Merox unit, and 

• retrofitting a UOP MTBE unit. 

In both cases, there is potential to significantly reduce the investment cost associated with FAME 

production, not only within the battery limits of the unit but also by using existing refinery 

infrastructure to minimize the overall costs. 

Merox Unit Retrofit 

UOP Merox (mercaptan oxidation) units have been used by refiners for over 50 years to remove 

mercaptan compounds from light fuel streams and to sweeten gasoline and jet fuel streams. In the 

Merox process, mercaptans are catalytically converted to non-odorous disulfide compounds. 

Depending on the application, one of several different processing schemes may be employed, 

involving mercaptan extraction or fixed-bed sweetening. Four alternatives are illustrated in Figure 

7.

Figure 7. UOP Merox processing schemes. 

When a new hydrotreater is installed to lower the amount of sulfur in a gasoline product, existing 

Merox equipment can become available for retrofit to bio-fuel service. For example, existing 

liquid extraction columns can be used as a feed or product wash column in a FAME production 

plant. Likewise, existing pumps, exchangers, reactors, separators, and caustic handling equipment 

can be reused in the esterification of soybean or rapeseed oil. 

9
 

Figure 8 illustrates a potential retrofit application wherein existing extraction and fixed bed 

conversion equipment is reused. The retrofit requires that a heterogeneous catalyst be developed 

for use in conjunction with an existing Merox fixed-bed reactor/separator to trans-esterify 

triglycerides in the presence of excess methanol and phase separate the crude glycerol product. 

Either a mixing device and/or co-solvent might be used to ensure adequate mixing of the 

triglyceride and methanol in the reactor, 

The FAME-rich product phase is then routed to an existing Merox extractor tower for further 

purification by counter-current water wash (and drying). The crude glycerol co-product could 

either be recovered and sold or routed to another refining unit (such as FCC) for conversion by 

cracking. Another option is to install new equipment for glycerol purification. Since glycerol 

prices are expected to drop significantly as world-wide biodiesel production rates increase, the 

reuse of crude glycerol as a refinery feedstock may be more economically attractive than 

purification. 

Bio-diesel

W

S

W

Bio-diesel to drier

Existing 

Extractive

Merox Unit

Existing 

Minalk FB

Merox Unit

Alcohol

Make-up

Glycerol

Recovery Glycerol

to water

recovery

Water

Make-up

Vent

Vegetable Oil

or Fatty Acid Excess Alcohol Recycle

Alcohol

Make-up

Bio-diesel

W

S

W

Bio-diesel to drier

Existing 

Extractive

Merox Unit

Existing 

Minalk FB

Merox Unit

Alcohol

Make-up

Glycerol

Recovery Glycerol

to water

recovery

Water

Make-up

Vent

Vegetable Oil

or Fatty Acid Excess Alcohol Recycle

Alcohol

Make-up

Bio-diesel

W

S

W

Bio-diesel to drier

Existing 

Extractive

Merox Unit

Existing 

Minalk FB

Merox Unit

Alcohol

Make-up

Glycerol

Recovery Glycerol

to water

recovery

Water

Make-up

Vent

Vegetable Oil

or Fatty Acid Excess Alcohol Recycle

Alcohol

Make-up

Figure 8. Possible Merox unit retrofit to FAME service. 

MTBE Unit Retrofit 

UOP MTBE units have been used by refiners for about 25 years to convert C4 olefins from fluid 

catalytic cracking units (FCCs) to methyl tertiary butyl ether (MTBE). The process is somewhat 

analogous to a FAME production unit in that methanol is reacted with the feedstock to convert it 

to a fuel-blending component. Figure 9 illustrates one version of the MTBE flow scheme. While 

many MTBE units are still operating in refineries elsewhere, units located in US refineries are 

being shut down or revamped to alternate service due to new regulations prohibiting the addition 

of ethers to gasoline. 

10
 

Figure 9. UOP MTBE (Ethermax) technology 

A typical MTBE unit includes one or more fixed bed reactors where methanol (or ethanol) is 

reacted with butenes to form the corresponding ether. A portion of the reactor effluent is recycled 

to control the reactor exotherm. Unreacted feed components are separated from the higher-boiling 

ether product in a product fractionator. In Ethermax units, catalytic packing is used in the 

rectification section of the fractionating column to enhance conversion. Excess alcohol is then 

recovered and recycled back to the primary etherification reactor. 

Retrofit of existing MTBE fixed-bed reactors to trans-esterify triglycerides in the presence of 

excess methanol also requires that a heterogeneous catalyst be developed. To ensure adequate 

mixing of the triglyceride and methanol in the reactor, either a pre-mixer and/or co-solvent would 

be used. The existing fractionation section would have to be modified to separate the higher 

boiling point FAME and glycerol products. However, it is expected that existing alcohol recovery 

equipment could be reused in biodiesel service. 

III.1.2 Retrofit Enablers 

The economic viability of retrofitting existing equipment for FAME production from 

biofeedstocks is highly dependent upon identifying a heterogeneous catalyst that can achieve high 

single-pass conversion under the normal reactor operating conditions of existing units (<1000C). 

Because a typical FAME production unit has a relatively low processing capacity compared to 

most existing Merox or MTBE units, the unit turndown needs to be evaluated on a case-by-case 

basis. In principle, excess processing capacity can be judiciously used to provide longer reaction 

and/or separation times and to enhance high one-pass conversions of triglycerides. 

Preliminary Heterogeneous Catalyst Study 

There are incentives for development of a heterogeneous catalyst for FAME production whether 

it is employed in an existing or new unit: 

• reduce operating costs (specifically consumption of chemicals) 

• simplify product purification (eliminate caustic, acid, and salt byproducts) 

• reduce generation of unwanted byproducts (soap stocks, wastewater, etc.) 

• produce a more valuable glycerol product 

Preliminary economic analyses suggest that cash cost of FAME production can be reduced by 

~5% or more with the development of a heterogeneous catalyst. 

11
 

A series of small-scale batch autoclave tests was performed to screen potential catalyst candidates 

using refined soy bean oil feedstock purchased from Aldrich and methanol. For each experiment, 

methanol, vegetable oil and catalyst were added to the autoclave in a prescribed ratio after which 

the vessel was sealed, heated to reaction temperature and stirred for a prescribed amount of time. 

In some cases, a co-solvent was added to ensure good contacting between the liquid phases. 

Afterwards, the products were phase-separated, excess methanol and co-solvent was removed and 

the oil phase was sent for analyses. A summary of operating conditions and analyses of the oil 

product from these tests is summarized in Table 2. 

Table 2. Summary of FAME experimental results. 

Several catalyst formulations (M1 and B4-B8) showed high one pass conversion of triglyceride to 

FAME at mild operating conditions (<1000

C). These catalysts all require further development to 

ensure that the high activity can be maintained under the expected continuous flow conditions 

associated with a commercial processing environment. 

<---Glyceride Compounds --->

T & t Run CoSolvent Catalyst FAME MONO DI TRI Glycerol

55C, 2 hrs 486 Yes A1 0.18 0.01 0.68 99.13 0.01

55C, 2 hrs 487 Yes A1 0.05 <0.01 1.09 98.86 <0.01

80 C. 4 hrs 496 Yes A2 0.08 0.01 0.77 99.14 <0.01

80 C. 4 hrs 496 Yes A3 0.08 0.01 0.77 99.14 <0.01

55 C. 2 hrs 481 Yes I1 0.6 0.04 1.62 97.74 0.00

55 C. 2 hrs 482 Yes I2 0.01 0.10 1.73 98.16 <0.01

80 C. 4 hrs 497 Yes I3 0.65 0.04 1.89 97.42 <0.01

80 C. 4 hrs 495 Yes B1 3.08 0.36 6.44 90.12 <0.01

55C, 2 hrs 489 Yes C1 6.58 1.09 12.43 79.87 0.03

55C, 2 hrs 493 No C2 22.82 0.51 6.73 69.87 0.07

55C, 2 hrs 484 No M1 54.73 0.79 6.37 37.62 0.49

55C, 2 hrs 485 Yes M1 77.28 6.74 6.96 8.72 0.29

80C, 4 hrs 490 Yes M1 86.32 5.18 4.45 3.45 0.60

80C, 4hrs 491 Yes M1 10.63 1.01 9.96 78.36 0.04

55C, 4 hrs 492 Yes AM1 31.44 0.53 4.23 63.53 0.27

60C, 4 hrs 498 Yes AM1 31.97 0.27 3.43 64.1 0.23

Solid <---Glyceride Compounds --->

T & t Run CoSolvent Catalyst FAME MONO DI TRI Glycerol

80 C, 4 hrs 500 Yes B2 0.02 0.01 0.67 99.31 0.01

80 C, 4 hrs 501 Yes B3 2.61 0.14 9.0 88.25 0.01

80 C, 4 hrs 502 Yes B4 97.78 1.35 0.72 0.06 0.09

80 C, 4 hrs 503 Yes B5 97.45 1.62 0.67 0.01 0.26

80 C, 4 hrs 509 No B5 97.48 0.88 0.6 0.04 1.04

80 C, 4 hrs 511 Yes B6 96.70 2.33 0.86 0.01 0.11

80 C, 4 hrs 506 Yes B7 81.63 8.45 8.31 1.07 0.54

80 C, 4 hrs 507 Yes B8 73.71 5.22 14.91 5.75 0.41

Financial Incentives For Retrofits 

There are additional financial incentives associated with retrofitting existing refining equipment 

such as Merox and MTBE units to FAME production, including: 

12
 

• leveraging existing land and refinery infrastructure, 

• reducing capital investment for processing equipment, and 

• capturing subsidies associated with production of renewable fuels. 

However, the metallurgy of the existing equipment (typically carbon steel) may limit the 

processing of high free fatty acid (FFA) feedstocks. Preliminary economic analyses suggest that it 

is possible to reduce the overall investment requirements by 50% or more. Coupling the lower 

capital expenditures with renewable fuel subsidies make the potential payback on new capital 

reasonably short (<2 years). 

III.1.3 Conclusions 

In this preliminary analysis, potential retrofits of existing Merox and MTBE units to FAME 

production appear promising under the right circumstances: 

• a heterogeneous catalyst is developed for operation under mild conditions (<1000

C) 

• an appropriate combination of feed price, feed availability, and product market. 

• adequate metallurgy for the proposed feedstock in the existing unit 

Several promising heterogeneous catalysts were identified in UOP’s laboratory screening test. 

However, the screening test only evaluates initial catalyst activity. The next step is to subject the 

most promising candidates to continuous-flow testing at operating conditions consistent with the 

anticipated commercial processing environment to ensure the high activity can be maintained. 

It should be noted that for all the biodiesel catalysts that achieved high conversion, it was 

discovered that base had leached into the liquid phase so these tests were not completely 

heterogeneous but had a homogenous component as well. When the leachable base was removed 

from the catalyst with highest conversion, the conversion was much lower (less than 50%). 

Therefore further work to obtain a stable base catalyst will be needed in order to commercialize 

this process. 

Assuming that an acceptable catalyst formulation is identified, then a realistic set of assumptions 

for plant capacity, feedstock cost, feedstock quality and product requirements need to be 

identified as the basis for a more detailed feasibility study. 

13
 

III.2 GREEN DIESEL PRODUCTION

“Green diesel” refers to an acceptable diesel pool blend component produced from a suitable 

biofeedstock. In contrast to biodiesel discussed in Section III.1, green diesel hydrocarbons do not 

contain oxygen in their molecular structure. Biofeedstocks that are of particular interest for green 

diesel production are plant oils and greases. 

III.2.1 Green Diesel Production Approaches 

Two alternatives were evaluated for producing green diesel from biofeedstocks, illustrated in 

Figure 10: 

• (a) co-processing oil or grease in an existing hydroprocessing unit, and 

• (b) processing oil or grease in a standalone process unit. 

The two case studies are summarized in Sections III.2.2 and III.2.3, respectively. Supporting 

experiments are summarized in Section III.2.4. 

III.2.2 Green Diesel Production – Co-Processing Biofeedstocks in an Existing Unit 

(a) Pretreat and Co-process

Distillate 

Hydrotreater

Pretreater

Diesel

H2

H2

Oil or grease

(b) Stand Alone Skid Mounted Unit

Oil or grease

Diesel H2

Green Diesel

H2

Diesel 

(ULSD) 

Distillate 

Hydrotreater

Diesel 

(ULSD) 

(a) Pretreat and Co-process

Distillate 

Hydrotreater

Pretreater

Diesel

H2

H2

Oil or grease

(b) Stand Alone Skid Mounted Unit

Oil or grease

Diesel H2

Green Diesel

H2

Diesel 

(ULSD) 

Distillate 

Hydrotreater

(a) Pretreat and Co-process

Distillate 

Hydrotreater

Pretreater

Diesel

H2

H2

Oil or grease

(b) Stand Alone Skid Mounted Unit

Oil or grease

Diesel H2

Green Diesel

H2

Diesel 

(ULSD) 

Distillate 

Hydrotreater

Diesel 

(ULSD) 

The process unit for the biofeedstock co-processing a biomass source was a distillate hydrotreater 

for ultra-low sulfur diesel (ULSD). The base case was a hydrotreating unit with the following 

specifications: 

Figure 10. Alternatives for green diesel production. 

• straight-run diesel feed from a blend of Middle-East crudes with less than 10% FCC light 

cycle oil 

• 35,000 bpd, 950 psig separator pressure, 1.1 LHSV, and 6,000 scfb total gas rate 

• chemical hydrogen consumption of 1.1wt.% relative to fresh feed 

• heat release of 160 Btu/lb 

• 30 month cycle time with a state-of-the-art Co/Mo catalyst. 

• activity limited catalyst cycle with a catalyst temperature cycle of 27°F. 

14
 

The objective of the co-processing case study was to 

produce a “G5” diesel pool in which 5% of the diesel pool 

contains “green diesel”. Since the United States does not 

currently subsidize bio-diesel or green diesel production, 

the biomass source of choice is one with an inexpensive 

feed cost, such as brown grease with 100% free fatty 

acids. The hydroprocessing conditions are assumed to 

completely favor hydrodeoxygenation (HDO). In other 

words, oxygen in the biofeedstock is removed by reaction 

with hydrogen to form water and the C18 n-paraffin. For 

G5 diesel production at constant crude-based diesel feed 

rate, the biomass feed rate to the distillate hydrotreating 

unit must be approximately 5.8% of the crude feed rate, or 

2050 bpd relative to the base case. In addition, the crude 

rate and crude slate are not changed to accompany the 

additional feed into the distillate hydrotreating unit. The 

estimated distillate yield is shown in Table 3. These yields 

assume no yield loss from production of light ends such as 

methane, ethane, LPG components, or naphtha. 

Table 3. Hydrotreating yields 

for brown grease. 

Components Wt% 

Feeds 

Brown grease 100 

H2 2.9 

Products 

H2S + NH3 0.04 

C1 to C4 0.7 

Naptha 2.4 

CO2 0.0 

H2O 12.7 

nC17 0.0 

nC18 87.2 

Total 102.9 

Table 4 summarizes the expected hydrogen consumption and heat release for fatty acid 

components of the model brown grease, compared with the crude oil-based feed. Using the model 

brown grease composition shown in the third column, the estimated hydrogen consumption for 

hydrodeoxygenation is 3.0 wt%. The anticipated heat release is approximately 385 Btu/lb. 

Table 4. Hydrotreating free fatty acids. 

Feed C:= 

% in 

Model 

Feed MW 

n H2 / 

FFA 

H2 

Consumed 

(wt% feed) 

Total Heat 

Release 

(Btu/lb) Br No 

Free Fatty Acid 

Stearic Acid 18:0 25 284.47 3 2.13 175 0 

Oleic Acid 18:1 45 282.46 4 2.85 367 57 

Linoleic Acid 18:2 25 280.44 5 3.59 562 114 

Linolenic Acid 18:3 5 278.43 6 4.34 760 172 

Diesel 1.08 160 2.5 

The effects of the following factors on the start-of-run average bed temperature (activity) and 

deactivation rate were analyzed using UOP-proprietary distillate hydrotreating reactor 

correlations: 

• properties of the brown grease feed such as sulfur, oxygen, and olefin content 

• heat release from olefin saturation and hydrodeoxygenation of brown grease 

• increased space velocity 

• “treat-gas” ratio, or the total gas (makeup plus recycle gas) ratio 

Three cases were analyzed to demonstrate the individual effects of the variables relative to the 

base case: 

15
 

• Case I demonstrates the effect of brown grease feed properties. Case I is the base case 

with 5.8% biomass added to the fresh feed while maintaining the total fresh feed rate. 

The crude-based feed rate is reduced to approximately 33100 bpd and supplemented with 

1900 bpd brown grease. The product contains approximately 5% green diesel, as desired. 

• Case II demonstrates the effect of increasing space velocity to accompany the additional 

biomass feed. Case II maintains the ratio between the feed types of Case I and increases 

total fresh feed rate to maintain the feed rate of crude in the base case. The refinery crude 

slate and crude rate is not changed from the base case. The treat-gas ratio is maintained at 

the base-case ratio, under the assumption that the recycle gas compressor is not capacitylimited. 

• Case III builds upon Case II by limiting the recycle gas compressor capacity to the basecase absolute capacity. 

The results of analyzing the three cases of brown grease co-processing are depicted in Figure 11

through Figure 13; the results are expected to represent the general behavior of low or no-sulfur 

biofeedstocks such as plant oils and greases. 

Figure 11 shows the effect of co-processing 

brown grease on the catalyst temperature cycle, 

defined as the end-of-run (EOR) peak 

temperature minus the SOR peak temperature. 

For the purpose of this study, the end-of-run 

peak temperature is fixed at the base-case value. 

Brown grease shows has an activity advantage 

regarding desulfurization. In this study, the 

start-of-run (SOR) average bed temperature 

(ABT) is reduced by approximately 2°F. 

However, the activity advantage is eroded by 

the heat release of brown grease, which is 

nearly 2-1/2 times higher than diesel. When the 

space velocity is increased in Case II, the 

corresponding increase in released heat raises 

the SOR ABT by approximately 3°F, consequently decreasing the catalyst temperature cycle. 

Capping the recycle gas compressor absolute capacity in Case III limits the system’s ability to 

control the catalyst bed temperature rise and, hence, the catalyst SOR ABT increases further, 

decreasing the temperature cycle by almost 4°F relative to the base case temperature cycle of 

27°F. 

-5

-4

-3

-2

-1

0

1

 T (°F)

Base Case I Case II Case III

from WABT T Cycle Difference

Figure 11. Effect of co-processing brown 

grease on catalyst temperature cycle. 

0.90

0.95

1.00

1.05

1.10

1.15

R elative D e activation 

Rate (Base = 1.00)

Base Case I Case II Case III

Figure 12 shows the relative deactivation 

rates for the cases analyzed. The sulfur 

content of the overall feed decreases when 

the crude-based feed is diluted with a 

biofeedstock containing little or no sulfur. 

The severity of the hydrotreating catalyst 

processing conditions is consequently 

reduced for the same product sulfur content. 

As a result, the deactivation rate decreases. 

As a consequence of further increasing the 

feed rate and decreasing the hydrogen 

circulation rate over the catalyst, the catalyst 

deactivation increases by approximately 

Figure 12. Effect of co-processing brown 

grease on catalyst deactivation rate. 

16
 

20% over Case I. One important factor not included in this analysis of catalyst stability is water 

production. 

22

24

26

28

30

32

Nom inal Cycle Length (m o.)

Base Case I Case II Case III

The catalyst cycle length, shown in Figure 

13, is determined by the temperature cycle 

and the deactivation rate. The cycle length 

for G5 diesel production, without 

decreasing the crude-based diesel rate or 

increasing the recycle gas compressor 

capacity, decreases by 6 to 7 months (Cases 

II and III). This reduced cycle time has 

economic consequences - for example, 2 to 

3 days per year production loss and 

increased catalyst and associated turnaround 

costs. 

The study also evaluated the impact of co-processing brown grease on reactor metallurgy and on 

the cloud point of the resulting diesel pool. The feed section metallurgy is carbon steel in 

hydroprocessing units such as hydrotreaters and hydrocrackers. There is a graduation of 

metallurgy from low-chrome through austenitic stainless steel in the combined feed exchangers. 

The fired heater tubes, reactor cladding and associated reactor circuit piping are typically 321 or 

347 stainless steels. The Total Acid Numbers (TAN) of plant oils and greases range from 2 to 

200. A TAN number of 0.5 to 0.6 is an industrially-accepted limit before considering upgrades to 

the process unit metallurgy. Therefore, the reactor section metallurgy of a hydroprocessing unit 

must be upgraded if any significant quantity of an economically-attractive biofeedstock is to be 

processed. The cost was estimated for upgrading the reactor section metallurgy to 317L stainless 

steel in the base case distillate hydrotreater and found to be substantial – approximately 7 to 8 

million dollars. Additional capital costs that were not included are costs for brown grease tankage 

and brown grease pretreatment equipment for possible dewatering, acid washing, ion exchange, 

and/or desalting. 

Figure 13. Effect of co-processing brown 

grease on catalyst cycle length. 

Distillates from brown grease and other similar biofeedstocks contain primarily n-paraffin. The nparaffin content will decrease the cloud point of the diesel pool from 2 to 3°F per percent of nparaffin content. Therefore, assuming that all diesel components in the diesel pool are 

hydrotreated to make ultra-low sulfur diesel, G5 diesel production may increase the cloud point 

of the diesel pool by 10 to 15°F. A diesel isomerization catalyst would be required to counteract 

the cloud point increase and preserve the green diesel molecules in the diesel pool. Unfortunately, 

such isomerization catalysts are typically noble metal catalysts and would require a two-stage 

hydroprocessing scheme, which is not typical for distillate hydrotreating units, necessitating 

major unit modifications. 

In summary, co-processing of biomasses such as plant oils and greases in a typical refinery 

distillate hydrotreating unit has some drawbacks for green diesel production: 

• The chemical nature of the biofeedstocks significantly increases hydrogen requirements 

under conditions that favor hydrodeoxygenation. 

• The heat release from these feeds substantially reduces the catalyst cycle length. 

• High TAN number feeds incur significant capital costs for metallurgy upgrading. Other 

capital costs may include feed pre-processing equipment to remove contaminants. 

17
 

• Typical hydrotreating unit configurations do not permit remediation of increased cloud 

point due to n-paraffins in the green diesel component of the diesel pool. 

• The process conditions for hydrotreating units for ultra-low sulfur diesel production favor 

HDO, producing water that may adversely affect catalyst performance. 

In conclusion, standard refinery distillate hydrotreating units do not appear to be suitable for 

green diesel production in a co-processing scheme. 

III.2.3 Green Diesel Production – Processing Biofeedstocks in a Stand-Alone Unit 

Green diesel production in a standalone unit appears to address the issues identified in Section 

III.2.2 for co-processing schemes: 

• catalyst performance can be improved through feed pretreatment 

• a process can be designed to favor a different deoxygenation mechanism 

• a simpler isomerization process can be used to prevent cloud point reduction 

• the unit can be specifically designed to handle biofeedstocks 

One possible flow scheme for a standalone unit is shown in Figure 14. 

Heater

Reactor

Separator

Stripper

Tall Oil

Feed

Recycle Gas

Compressor

Light Ends

Diesel

Product

Hydrogen

Makeup

Makeup

Compressor Sour

Water

PSA CO2

CH4 H2 Plant

Feed 

pump

Cooler

Heater

Reactor

Separator

Stripper

Tall Oil

Feed

Recycle Gas

Compressor

Light Ends

Diesel

Product

Hydrogen

Makeup

Makeup

Compressor Sour

Water

PSA CO2

CH4 H2 Plant

Feed 

pump

Cooler

Plant oils and yellow and brown greases may contain water, alkali metals (such as sodium, 

potassium), phosphorous, ash, and other contaminants. Metals permanently deactivate 

hydroprocessing catalysts. In a standalone unit, a pre-processing step would remove contaminants 

from the biofeedstock. using a combination of technologies such as hydrocyclones, desalting, 

acid washing, ion exchange, or fixed guard bed catalyst systems. This is a distinct advantage over 

Figure 14. Proposed flow scheme for standalone production of green diesel. 

18
 

a co-processing system. A typical strategy for metals removal is to provide a less active sacrificial 

bed of hydrodemetallization (HDM) catalyst in front of the active hydrodesulfurization (HDS) 

catalyst. In existing distillate hydrotreaters, either active catalyst would be removed to load HDM 

catalyst or the metals would be allowed to deposit on the active HDS catalyst. Both choices 

would decrease the catalyst cycle length of an existing unit. 

The processing conditions to produce ultra-low sulfur 

diesel (ULSD) in an existing unit highly favor HDO, 

with the drawbacks outlined in Section III.2.2. An 

alternative mechanism, decarboxylation (DeCO2), 

removes oxygen in the biomass carboxyl groups as 

carbon dioxide instead of as water. In a standalone unit, 

catalyst and processing conditions can be selected to 

shift the oxygen removal mechanism from HDO to 

DeCO2. Decarboxylation and hydrodeoxygenation are 

illustrated in Figure 15. 

Cn COOH

Pd / H

Cn + CO2

Cn COOH Cn+1 + 2H2O

+3H2

Decarboxylation

Hydrodeoxygenation

Cn COOH

Pd / H

Cn + CO2

Cn COOH Cn+1 + 2H2O

+3H2

Decarboxylation

Hydrodeoxygenation

Figure 15. Decarboxylation vs 

Hydrodeoxgenation 

Gas phase DeCO2 is discussed for several model 

compounds by Maier.12 Table 5 compares the 

estimated yields of DeCO2 with HDO for the 

model brown grease. There are several clear 

advantages for DeCO2: 

Table 5. Hydrodeoxygenation versus 

decarboxylation theoretical yields 

Components 

DeCO2 

Wt% 

HDO 

Wt% 

Feeds 

Brown grease 100 100 

H2 0.8 2.9 

Products 

H2S + NH3 0.04 0.04 

C1 to C4 0.6 0.7 

Naptha 2.5 2.4 

CO2 15.5 0.0 

H2O 0 12.7 

nC17 82.2 0.0 

nC18 87.2 87.2 

Total 100.8 102.9 

• Chemical hydrogen consumption is 

limited to olefin saturation, therefore 

significantly lower. 

• Lower chemical hydrogen consumption 

reduces the hydrogen plant capital or 

hydrogen purchases and the makeup 

hydrogen compressor size. 

• Water is not produced as a reaction 

product, thereby preserving catalyst 

performance. 

• Rejected CO2 can be captured in a 

relatively pure state as a product. From a 

greenhouse gas emissions viewpoint, the 

proposed green diesel process would not 

only be CO2 neutral, but actually produce 

additional CO2 credits. 

Decarboxylation can not be considered in a unit producing ULSD for two reasons: (1) DeCO2 is 

favored over HDO at lower temperatures and pressures than the conditions required for deep 

desulfurization of traditional feedstocks; (2) the production of CO2 will reduce hydrogen partial 

pressure unless the purge rate from the high-pressure separator, and subsequently, the make-up 

hydrogen rate are both increased. In a stand-alone unit, a pressure swing adsorption unit (PSA) 

system in the recycle gas stream can be provided to remove CO2. In this way, recycle gas purge, 

recycle gas and makeup gas compression are minimized. 

The relevant biofeedstocks have low concentrations of sulfur and nitrogen compounds relative to 

standard feedstocks. This makes n-paraffins in the green diesel product amenable to 

hydroisomerization without a true two-stage hydroprocessing unit such as would be required in a 

19
 

co-processing scheme. A hydroisomerization reactor directly follows the olefin saturation / 

decarboxylation stage in the flow scheme of Figure 14. 

A standalone unit can be design from the outset with metallurgy appropriate to the high TAN 

numbers of the candidate biofeedstocks. Upgrading most of the metallurgy to 317L stainless steel 

allows poorer quality and less expensive feeds (higher free fatty acid content) to be processed 

than would be possible in a typical hydroprocessing unit. However, the capital cost of the unit is 

obviously increased. Coupon testing must be done to determine whether the proposed metallurgy 

provides adequate corrosion protection. The estimated inside battery limits capital cost for a 2000 

bpd standalone green diesel unit is approximately $6,000 to 7,000 per barrel. 

III.2.4 Green Diesel Experimental Results 

Background information on green diesel is available from Arbokem/Canmet, Fortum Oil, and 

most recently OMV. Arbokem has the key patent in green diesel production which expires in 

2008.13 Initially it was thought that yield estimates could be generated from literature and 

modeling results alone. Our initial yield estimates assumed that hydrodeoxygenation would be the 

primary mechanism for oxygen removal, as reported by Canmet/Arbokem from their initial 

studies14 run at 700-2200 psi and 350-450C. 

However, experiments performed in this project showed that both hydrodeoxygenation and 

decarboxylation occur, with the results summarized in Table 6. Decarboxylation is favored at 

lower pressures and hydrodeoxygenation increased with increasing pressure. Decarboxylation 

results in odd number paraffin production and CO2 formation whereas hydrodeoxygenation 

results in even carbon number paraffin production; therefore the ratio of nC17 to nC18 is a 

measure of the DeCO2/HDO ratio. Standard hydrotreating catalysts of NiMo, CoMo and Pd all 

showed activity for both reactions. Furthermore it was demonstrated that the process could be run 

at lower temperatures than cited in the Arbokem patent and heavy material boiling above diesel 

range, which Arbokem produced, could be eliminated or reduced with proper choice of catalyst 

and conditions. Also, most experiments were done at 500psi which is well below the pressures 

cited in the Arbokem patent. 

Table 6. Green diesel experimental summary. 

524 531 1316-C3 1316-C4 

Feed Vegetable oil Vegetable oil Vegetable oil Vegetable oil 

Catalyst HCT (NiMo) N200 (CoMo) HCT (NiMo) HCT (NiMo) 

Test Unit Autoclave Autoclave Plt 1316-B1 Plt 313-B1 

WHSV 1.9 1.7 0.8 0.3 

Temperature (C) 300-350 300-350 325 310 

Pressure (psia) 500 500 500 500 

Products 

% water 1.7 1.2 4.9 9.9 

% CO2 + CO 12.7 13.4 2.6 1.9 

% Light HC 7.0 5.2 3.4 3.9 

% diesel+ 79 80 88.3 84 

%heavy 0 3.2 0.3 0.3 

%deoxygenation 90+ 91+ 85 99 

nC17/nC18 3.0 2.1 0.6 0.49 

20
 

Further testing is recommended to develop an optimized process. These experiments were proofof-principle (POP) experiments, so good material balances were not achieved. In the autoclave, 

liquid and catalyst are charged to the reactor and H2 gas is constantly flowing. For autoclave tests, 

losses are treated as gas which is calculated by difference so gas yield is overestimated and liquid 

yield underestimated. Gas composition changed throughout the test period because the test was 

not at steady state and true material balance and gas weights were not obtained. 

The most active hydrotreating catalysts, such as UF-310 and UF-210 Stars, were not tested in 

these POP tests because of difficulties in presulfiding these materials. Therefore, there is 

considerable opportunity for improvement during development of a green diesel process, if 

pursued. 

III.3 COMPARISON OF BIODIESEL AND GREEN DIESEL PROCESS AND PROPERTIES

It is useful to compare biodiesel and green 

diesel properties and process since they 

represent two different approaches for 

making diesel fuel from vegetable oil. 

Selected properties are compared in Table 

7. Green diesel has excellent diesel fuel 

properties including an extremely high 

cetane number (see Section III.3). Cetane 

number is a measure of a diesel fuel’s 

ignition delay, e.g. the time period 

between the start of injection and start of 

combustion of the fuel, and affects the 

price of both green diesel (80 cetane) and 

biodiesel (55 cetane). Biodiesel has lower 

heating value because of its oxygen 

content. Biodiesel also has other 

undesirable properties such as high 

density, and high NOx emissions. Overall, 

green diesel appears to be a superior 

product. 

Table 7. Biodiesel and green diesel properties 

Biodiesel 

(FAME) 

Green 

Diesel 

%O 11 0

Density g/ml 0.883 0.78

Sulfur content <10ppm <10ppm

Heating Value 

(lower) MJ/kg 38 44

% change in 

NOx emission +10 0 to -10

Cloud Point C -5 -5 to -30

Distillation 

10-90% pt 340-355 265-320

Cetane 50 80-90

Table 8 compares the processes by which biodiesel 

and green diesel are produced. The biodiesel process 

is fairly complicated, requiring multiple distillations 

and purifications. Green diesel is totally feed 

flexible and can handle low cost feedstocks like 

brown grease with high free fatty acid content 

without problems. Biodiesel requires methanol as 

feed and produces glycerol as a byproduct. Glycerol 

price is expected to drop significantly to $.10/lb as 

biodiesel production increases. Green diesel only 

requires hydrogen as a feed and produces only diesel 

and propane as a product so its value is not affected 

by methanol or glycerol price. Biodiesel requires a 

homogeneous catalyst, resulting in higher chemical 

costs. 

Table 8. Comparison of biodiesel and 

green diesel production processes. 

Biodiesel Green Diesel 

Complexity Highmultistages low 

Feedstocks Sensitive to 

level of FFA 

No sensitivity 

to FFA 

Requires Methanol -

10% H2 -2-3% 

Byproduct Glycerol Propane 

Catalyst Homogeneous Heterogeneous 

21
 

As shown in Table 9, biodiesel does 

have an advantage in yields of 

diesel product per lb of vegetable oil 

feed. This is due to the fact that the 

oxygen in vegetable oil is retained 

in the biodiesel product where as the 

oxygen in vegetable oil is rejected 

green diesel as water or CO2. 

However, the overall cost of 

biodiesel production is projected to 

be nearly twice that of green diesel. 

Table 9. Green diesel and biodiesel yields. 

Feed Biodiesel Green Diesel 

% Oil or Grease 100 100 

% H2 1.5-3.8 

% methanol 8.7 

Products 

% water,CO2 12-16 

% Lt HC 2-5 

% diesel 96 83-86 

% glycerol 12 

Operating cost $/gal .05 .025 

III.4 GREEN GASOLINE AND 

GREEN OLEFINS 

PRODUCTION 

Vegetable oils and greases can be processed in an FCC to produce high yields of gasoline and 

oils. ACE microactivity tests were completed to determine the yields and product properties. The 

gasoline produced has high octane number and the vegetable oil produces a lot of high valued 

olefins at PetroFCC conditions (severe conditions design to maximize olefins production) 

compared to vacuum gas Oil (VGO), as shown in Table 10. The yields from processing vegetable 

oil and grease compared to VGO are shown in Table 10 and Table 11 . 

Table 10. Green olefin yields. Table 11. Green gasoline yields 

VGO Vegetable 

Oil/Grease 

C2P+Methane 4.1 4.1 

C2= 8.6 8.7 

C3P 2.0 2.1 

C3= 22.0 22.4 

C4s 15.0 13.5 

Gasoline 27.3 23.0 

LCO 9.5 5.0 

CSO 5.0 3.0 

Coke 6.5 6.5 

Water 0 11.7 

RON of Gasoline 94.8 96.8 

VGO 

Vegetable 

Oil/Grease 

C2= 1.5 1.9 

C3P 0.7 0.8 

C3= 4.0 4.6 

C4s 7.9 6.6 

Gasoline 45.5 45.4 

LCO 17.5 11.4 

CSO 19.5 13.1 

Coke 3.4 4.5 

Water 0 11.7 

RON of Gasoline 92.1 94.8 

The acid number for typical total FCC product is 0.016. In one experiment, the acid number of 

the biofeedstock was 0.6 which is slightly elevated and raises concerns about corrosion. At a 5% 

level, however, the biofeedstock contribution would be only 0.03, and the net acid number would 

be 0.045, which is well below the problem where corrosion issues occur. 

22
 

VGO

FCC

Pretreat

Gasoline

LCO

Lt Ends

CSO

VGO

FCC

Pretreat

Gasoline

LCO

Lt Ends

CSO

Figure 16. FCC processing approach. 

In order to process vegetable oil and grease a separate feed system and pretreater would be best to 

remove metals (e,g, Ca,K ) which would poison the FCC catalyst and to avoid metallurgy issues 

in the feed system when processing greases. A suggested processing scheme is shown in Figure 

16. 

III.5 ECONOMIC ASSESSMENT AND LIFE CYCLE ANALYSIS

Economic comparisons for the various 

processing options are shown in Figure 17

and Figure 18. The cost estimate procedure 

has the following qualifications: These costs 

represent U.S.A. Gulf Coast erection to UOP 

standards for new equipment on a January 

2005, open shop (non-union) labor basis. The 

equipment costs have an anticipated accuracy 

of -15%/+30%. The estimated erected cost 

(EEC) is a factored cost, which includes 

installation, associated bulk items (such as 

instruments, electrical, piping, and civil), 

indirects, and contractor’s home office 

expenses. These costs are classified as inside 

battery limits (ISBL) for this estimate. The 

EEC is derived by multiplying the total equipment cost by a factor—the factor being obtained 

from an historical database which records average equipment cost versus total erected cost from 

previous detailed cost estimates developed by UOP Cost Estimating. The EEC has an anticipated 

accuracy of -30%/+50%. The EEC is for battery 

limits only. Indirect costs, home office expenses 

included in the EEC are not based on single 

equipment item installations; these equipment 

items are assumed to be part of a larger / typical 

refinery project. 

0

10

20

30

40

Biodiesel Green Diesel Green Gasoline

$M

Stand alone

Pretreat + Coprocess

0

10

20

30

40

Biodiesel Green Diesel Green Gasoline

$M

Stand alone

Pretreat + Coprocess

Biodiesel Green

Diesel

Green

Gasoline

10

20

30

40

0

10

20

30

40

Biodiesel Green Diesel Green Gasoline

$M

Stand alone

Pretreat + Coprocess

0

10

20

30

40

Biodiesel Green Diesel Green Gasoline

$M

Stand alone

Pretreat + Coprocess

Biodiesel Green

Diesel

Green

Gasoline

10

20

30

40

Figure 17. Capital costs of biofuels 

production from oils and greases. 

-300

-200

-100

0

100

200

$M

Soy-Unsubsidized

Soy-subsidized

Yellow Grease

Brown Grease

-300

-200

-100

0

100

200

$M

Soy-Unsubsidized

Soy-subsidized

Yellow Grease

Brown Grease

Biodiesel Green 

Diesel

Green 

Gasoline

Green 

Olefins

-300

-200

-100

0

100

200

$M

Soy-Unsubsidized

Soy-subsidized

Yellow Grease

Brown Grease

-300

-200

-100

0

100

200

$M

Soy-Unsubsidized

Soy-subsidized

Yellow Grease

Brown Grease

Biodiesel Green 

Diesel

Green 

Gasoline

Green 

Olefins

Figure 18 summarizes the economic analysis of 

biofuels and chemicals production from oils and 

greases through comparison of their net present 

values (NPV). The NPVs of four products are 

shown as function of feedstock. Biodiesel is the 

least cost-competitive product for any feedstock, 

due to the high capital cost of biodiesel 

production. 

Figure 18. NPV comparison of biofuels and 

chemicals. 

23
 

The effects of potential subsidies on NPV are illustrated in Figure 18.

15, , , 16 17 18 Brown grease does 

not require subsidies to be an economically-viable feed for any of the products, because of its low 

cost. However, subsidies would have a significant impact on the viability of all soy-based 

products, as can be seen by comparing the NPV with and without subsidies. Note that only 

biodiesel is subsidized currently in the U.S, while all options are on equal footing in the EU. 

A life-cycle analysis (LCA) of the various vegetable oil processing routes was conducted using 

the Simapro LCA program. LCA is a method to determine and compare the environmental impact 

of alternative products or processes “from cradle to grave”. In this case, the scope of the analysis 

was from extraction through combustion. For analysis purposes, it was assumed that all fuels 

have the same performance in transportation use. In reality, biodiesel contains 12% oxygen and 

has much lower fuel efficiency than the other fuels; this effect is not reflected in the results. The 

primary focus of the analysis was on fossil energy consumption and emission of greenhouse 

gases, though other impact categories are included. 

The results of the analysis are shown in Figure 19 and Figure 20. In general,, the green products 

have much lower total environmental impact score (Figure 19) and produce less CO2 (Figure 20) 

than petroleum diesel. Of the biofuels, green diesel and green gasoline have the lowest 

environmental impacts and CO2 production. The environmental impacts of biodiesel production 

are increased by the fact that it requires methanol which is produced from natural gas. Methanol 

production is a very energy-intensive process so it brings with it a strong environmental burden. 

0

50

00

50

00

50

00

Petroleum

Diesel

Biodiesel Green

Diesel

Green

Gasoline

0

50

00

50

00

50

00

Petroleum

Diesel

Biodiesel Green

Diesel

Green

Gasoline

Points

0

50

00

50

00

50

00

Petroleum

Diesel

Biodiesel Green

Diesel

Green

Gasoline

0

50

00

50

00

50

00

Petroleum

Diesel

Biodiesel Green

Diesel

Green

Gasoline

Points

 

Figure 19. LCA – single environmental 

impact score . 

0

1

2

3

4

Petroleum

Diesel

Biodiesel Green

Diesel

Green

Gasoline

0

1

2

3

4

Petroleum

Diesel

Biodiesel Green

Diesel

Green

Gasoline

kg CO2/kg Fuel

0

1

2

3

4

Petroleum

Diesel

Biodiesel Green

Diesel

Green

Gasoline

0

1

2

3

4

Petroleum

Diesel

Biodiesel Green

Diesel

Green

Gasoline

kg CO2/kg Fuel

Figure 20. LCA –comparison of climate-active 

CO2 

24
 

IV REFINERY OPPORTUNITIES FOR PYROLYSIS OIL

Pyrolysis oil, or bio-oil, is the material produced from pyrolysis of wood waste. Pyrolysis oil can 

be used as a substitute for fossil fuels to generate heat, power, and/or chemicals. Transportation 

fuels such as methanol and Fischer-Tropsch fuels can be derived from pyrolysis oil through 

synthesis gas processes. Pyrolysis oil is also a potential feedstock for hydrogen production. In 

addition, a wide range of chemicals can be extracted or derived from pyrolysis oil. 

Section IV.1 outlines the properties of pyrolysis oil and alternatives for its utilization in refineries. 

This study investigated several processing options: 

• Hydrotreating pyrolytic lignin to produce aromatics and gasoline (Section IV.2) 

• Co-processing with vacuum gas oil (VGO) (Section IV.3) 

• Aqueous Phase Reforming (APR) of water-soluble components to produce hydrogen 

(Section IV.4) 

The study also investigated hydrothermal processing of pyrolysis oil and techniques for blending 

pyrolysis oil with other feedstocks (Section IV.5). 

IV.1 BACKGROUND AND PROCESSING OPTIONS

The properties of pyrolysis oil depend on the specific feedstock and conditions of the production 

process, such as temperature, period of heating, ambient conditions, and the presence of oxygen, 

water and other gases. A typical process for producing pyrolysis oil is shown in Figure 21. Wood 

biomass such as sawdust, forest “slash”, or wood waste from mills is ground to very small sizes 

and injected into a fluidized bed reactor. The fluidized bed typically consists of a circulating, noncatalytic material such as sand. The temperature and contact time are optimized to maximize 

liquid yields. Typical products from the pyrolysis process are gases, such as carbon oxides and 

light hydrocarbons like methane and ethane, and the pyrolysis oil itself, which consists of watersoluble oxygenated compounds and “pyrolytic lignin”. 

Exhaust Exhaust Gas

 Net Electricity

Recycle Carrier Gas

Process

Electricity

 Wood

 Bio-Oil

Combustion Char

Air

Ash

Feed Handling 

and Drying

Gas Clean-Up 

and Oil 

Recovery

Pyrolysis Steam Cycle

Char 

Combustion

Figure 21. Typical pyrolysis process. 

25
 

The cost of pyrolysis oil is significantly affected 

by the high capital cost of the production complex. 

The economics of pyrolysis oil utilization could be 

improved by reducing these capital costs. The 

typical quantity of oil produced from forest waste 

in a pyrolysis unit is shown in Table 12. 

Table 12. Material balance for 

pyrolysis oil production. 

Pyrolysis oil consists of an oil fraction, insoluble 

pyrolytic lignin, and a mixture of water-soluble 

components. The pyrolytic lignin is similar in 

structure to the lignin originating from the wood 

biomass, but has been reduced in molecular weight from the thermal conversion in the pyrolysis 

process. Pyrolytic lignin is much more hydrocarbon-like than the water-soluble pyrolysis oil since 

it has less oxygen. It is produced from the 

pyrolysis oil either by adding water to 

precipitate out the insoluble components 

or by gravity separation. Gravity 

separation yields more pyrolytic lignin 

and a higher water content. Pyrolytic 

lignin produced by gravity separation had 

a much high acidity than pyrolytic lignin 

produced by water precipitation. This 

suggests that acetic acid and other 

carboxylic acids are entrained in the 

pyrolytic lignin during gravity separation. 

The general properties of pyrolysis oil are 

summarized in Table 13. 

Feed 

Biomass (forest waste, tpd) 2400 

Product 

Pyrolysis Oil (tpd) 1440 

Pyrolysis Oil (bpd) 7500 

Capital Cost $88 MM 

Table 13. Pyrolysis oil properties. 

It is believed that the pyrolytic lignin is 

produced from thermal degradation of the 

lignin fraction of the biomass and 

therefore has essentially the same structure as 

lignin except a lower molecular weight. The 

structure of lignin is shown in Figure 22. It 

consists mostly of individual benzene rings 

connected via oxygen and short propane linkages. 

Breaking this structure at the oxygen linkages 

should yield aromatic gasoline and propane. 

Pyrolysis 

Oil* 

Pyrolytic 

Lignin 

Water 

Soluble 

Pyrolysis 

Oil 

Weight % 

MF

100 30.0 70.0

%C 44.7 64.2 36.3

%H 7.2 6.1 7.6

%N .2 .3 0.16

%O 47.9 29.7 55.7

Acid 

Number

70 - 150 30 - 150 70 - 150

Heating 

value 

LHV Btu/lb

6,560 10,200 5,000

The options for utilizing pyrolysis oil in refineries 

are affected by its high acid number, high water 

content, high oxygen content and high metal 

content, particularly potassium and calcium. 

Pyrolysis oil processing schemes for the refinery 

are shown in Figure 23. Figure 22. Structure of lignin. 

26
 

Pyrolysis oil

Reforming

Pyrolytic 

Lignin

H2

HCK

Gasoline + 

Diesel

CO2

Hydrotreater

Pre-reformer

LCO

CH4

Coprocessing in Existing Units

New Unit

Pyrolysis oil

Pyrolytic 

Lignin

H2

Gasoline 

and Diesel Hydrotreater/hck

FCC

VGO Gasoline + 

Diesel

Boiler or 

Furnace

Steam or heat

Hydrothermal/

pretreater

Pyrolysis oil Upgraded pyrolysis oil

Pyrolysis oil

Reforming

Pyrolytic 

Lignin

H2

HCK

Gasoline + 

Diesel

CO2

Hydrotreater

Pre-reformer

LCO

CH4

Coprocessing in Existing Units

New Unit

Pyrolysis oil

Pyrolytic 

Lignin

H2

Gasoline 

and Diesel Hydrotreater/hck

FCC

VGO Gasoline + 

Diesel

Boiler or 

Furnace

Steam or heat

Hydrothermal/

pretreater

Pyrolysis oil Upgraded pyrolysis oil

Figure 23. Pyrolysis oil processing routes. 

Metals can be removed with guard beds or ion exchange. Metals removal is required before 

processing because these materials will typically poison catalysts. The high water content and 

very high oxygen content make it difficult to blend the pyrolysis oil with common refinery 

intermediate streams such as VGO. However it was found that proper additives can be used to 

form stable microemulsions. This is discussed more thoroughly in Section IV.3. 

The most serious problem for pyrolysis oil processing is the high acid number of the pyrolysis 

oils which causes corrosion in standard refinery units. Pyrolysis oils can probably be processed 

using 317 stainless steel cladding, which is not standard in refinery units. Therefore pyrolysis oils 

require pre-processing in a 317 stainless steel system to reduce the acid number before processing 

in typical refinery units. 

IV.2 GASOLINE AND AROMATICS PRODUCTION FROM PYROLYTIC LIGNIN

In previous studies, whole pyrolysis oil was hydrotreated at high pressures (2000 – 2500 psi ) and 

low space velocities(0.1 – 0.2 LHSV). The resulting hydrotreated oil was then cracked in an FCC 

or hydrocracker to produce gasoline. At these high pressures and low space velocities, 

hydrodeoxygenation predominates. Large quantities of hydrogen are required to generate water 

during hydrodeoxygenation because of the high level of oxygen (46%) in pyrolysis oil. This 

approach is unlikely to be commercially viable because of the high hydrogen requirement and the 

high capital cost of the hydrotreatment step. 

Because of its molecular structure, removing oxygen from pyrolytic lignin naturally generates 

materials in the gasoline boiling range. Pyrolytic lignin has a lower oxygen content than pyrolysis 

oil and therefore requires less hydrogen. Furthermore, decarboxylation is increasingly favored at 

lower pressure, which also reduces hydrogen consumption. Note that the same type of metal 

catalysts (Ni,Co,Pd,Pt) catalyze both reactions. 

Two catalysts were tested for hydrotreating pyrolytic lignin: 

• PNNL hydrotreated Dynamotive pyrolytic lignin using a Pd coated carbide catalyst in a 

continuous pilot plant. 

27
 

• UOP hydrotreated Dynamotive pyrolytic lignin using a Ni-Mo catalyst in a batch 

autoclave. 

The results are shown in Table 14. 

Both tests were successful at reducing 

the acid number, boiling point and 

oxygen content of the oil. The UOP 

catalyst appeared to be more active for 

oxygen removal at lower pressure and 

higher space velocity. Further tests 

will be required to optimize the 

process conditions and test the long 

term stability and activity of the UOP 

HCT catalyst (a newer-generation 

hydrotreating catalyst). However these 

tests suggest that pyrolytic lignin can 

hydrotreated under mild conditions, 

and that more active catalysts can be 

developed for this application to 

enable operation at higher space velocities and 

lower pressures. 

Table 14. Results from hydrotreating pyrolytic 

lignin. 

PNNL HT UOP HT 

WHSV 0.52 1.0 

LHSV 0.22 0.68 

Catalyst Pd/C NiMo 

Pressure 1900 – 2000 1500 

Liquid yield % 55.6 40.8 

% Oxygen removal 69 93 

% Oxygen in product 19.5 5.9 

Acid number of product 34 15 

% Naphtha in liquid 30 60 

Table 15. Hydrotreating+hydrocracking 

pyrolytic lignin. 

PNNL HT + HCK 

LHSV, Step 1 0.22 

LHSV, Step 2 1.5 

Catalyst, Step 1 Pd/C 

Catalyst, Step 2 Ni/Mo 

Pressure, Step 1 1900 – 2000 

Pressure, Step 2 1500 

Liquid yield % 32.4 

% Oxygen removal 96 

% Naphtha in liquid 80 

The product from the PNNL hydrothermal test 

was then hydrocracked using a NiMo catalyst in a 

batch autoclave. The combination of hydrotreating 

and hydrocracking resulted in 96% oxygen 

removal and a very high level of conversion to 

gasoline, with the properties shown in Table 15. 

An even better process would use a more active 

catalyst in the hydrotreating step such as the HCT 

catalyst. It is likely that the aromatics production 

can be increased and of naphthene production 

reduced by further process and catalyst 

optimization. 

Estimates of the overall yield were developed 

for the overall two step process, based on the 

experimental results. These estimates, shown in 

Table 16, were used as a basis for economic 

calculations. 

Table 16. Yield estimate for 

hydrotreating/hydrocracking 

pyrolytic lignin. 

Feed Wt% bpd 

Pyrolytic lignin 100 2250 

H2 4 – 5 

Product 

Light ends 15 

Gasoline 30 1010 

Diesel 8 250 

Water, CO2 51 – 52 

Commercially, hydrotreating and hydrocracking 

would be done in a single process unit with two 

catalyst vessels and intermediate water removal. 

However, since gravity-separated pyrolytic 

lignin has such a high acidity, hydrotreatment in 

vessels clad with 317 stainless steel would be 

required before coprocessing in existing 

hydroprocessing or hydrocracking units. The 

industry standard for refinery vessels is that the 

28
 

total acid number of the blend must be less than 0.6. Therefore the TAN for the hydrotreated 

lignin product would have to drop below 6 for a 10% blend or below 12 for a 5% blend. This 

should be achievable with a more active catalyst in the first step. 

The total quantities of gasoline produced and pyrolytic lignin processed is relatively small by 

refinery standards. After hydrotreating, which requires 317 SS, most of the lignin conversion and 

upgrading has been performed, so it will generally be more cost-effective to perform the entire 

process in a single new skid-mounted refinery unit (Figure 24). There are a few cases in which 

refinery hydrotreating/hydrocracking units have already been designed with 317 because they are 

processing high naphthenic acid crudes. These units would be ideal for processing pyrolytic 

lignin but are relatively rare. 

Hydrotreating/ 

decarboxylation Hydrocracking

Water

Gasoline or 

aromatics

Pyrolytic 

lignin

H2 H2

Hydrotreating/ 

decarboxylation Hydrocracking

Water

Gasoline or 

aromatics

Pyrolytic 

lignin

H2 H2

Figure 24. Process flow for hydrotreating/hydrocracking of pyrolytic lignin. 

IV.3 COPROCESSING WITH VGO IN AN FCC

Small FCC pilot tests (ACE tests) were run to determine the crackability of pyrolysis oil blended 

with vacuum gas oil (VGO), pyrolytic lignin blended with VGO, and neat hydrotreated pyrolysis 

oil from PNNL. In the blends, the VGO serves as a hydrogen donor. Stable blends are needed for 

the pilot plant injector system although two separate feed systems and nozzles could be used in 

commercial units. Stable blends were prepared using Hypermer surfactants. Hypermer is a nonionic polymeric surfactant made by Uniqema.19

The results are shown in Table 17. Compared to VGO, the pyrolysis oil, pryrolytic lignin and 

hydrotreated pyrolytic lignin tend to form high levels of coke. On an incremental basis, the 

pyrolysis oil formed 84% coke+water+COx and the lignin formed 70% coke +water+CO2 from 

the biofeeds. On an incremental basis the pyrolysis oil formed 16% coke and the pyrolytic lignin 

formed 27% coke. This is consistent with results reported on FCC cracking of neat pyrolysis oils 

Table 17. Yields for FCC conversion of pyrolysis oil, pyrolytic lignin 

and HT pyrolysis oil. 

29

VGO 

20%Pyrolysis 

oil 

80%VGO 

blend) 

20%Pyrolytic 

Lignin 

80% VGO 

Blend 

HT Pyrolytic 

Lignin 

(PNNL) 

C2- 2.0 3.3 3.6 3.8 

C3P 1.2 2.1 2.4 .7 

C3= 5.9 6.1 6.3 2.6 

C4 11.1 13.5 14.3 2.7 

Gasoline 42.7 40.6 41.3 28.8 

LCO 14.8 9.1 9.7 15.6 

CSO 18.5 4.8 4.7 6.2 

coke 3.8 7.1 9.2 16.1 

Water + CO2 0 13.5 8.5 23.5 
 

and the carbon residue results for the samples shown in Table 18.

20, , 21 22

Table 18. Carbon residue results 

Carbon Residue Test 

Pyrolysis oil 21 

Pyrolytic lignin 27-29 

Hydrotreated pyrolysis oil (HT82) 6.7 

For the blends of VGO with pyrolysis oil or pyrolytic lignin, the acid bio-oils appeared to 

increase the crackability of the VGO and shift VGO yields toward increased light ends and lower 

LCO and CSO which is an economically attractive outcome. The neat hydrotreated lignin (HT82) had unexpectedly high coke levels. Any follow-up tests should test blends of the hydrotreated 

lignin with VGO to determine whether any synergies exist for blending this material with VGO. 

VGO and hydrotreated lignin form stable 

blends without need for surfactant. The 

high levels of coke obtained with 20% 

blends would be unacceptable for most 

FCC units but 5% blends might be 

possible. Further testing of 5% blends 

would be required to determine the 

viability of this approach. The properties 

of full boiling range liquid samples 

produced from the pilot plant test are 

shown in Table 19. 

Table 19. Product properties of total liquid 

products from FCC testing at 1000F and 5/1 

cat/oil ratio 

80% VGO and 

20% pyrolytic 

lignin 

Hydrotreated 

pyrolytic lignin 

%C 88.6 89.5 

%H 9.3 9.5 

%O .38 .56 

As an alternative to blending, coprocessing pyrolytic lignin and pyrolysis oil with VGO in an 

FCC might be possible if a separate feed system were used to inject the pyrolysis oil or pyrolytic 

lignin. There are three concerns which would need to be addressed in order to determine the 

viability of this approach: 

• Would the metallurgy of the FCC riser system be affected by injecting highly acid 

material or would the material react quickly enough to avoid problems? 

• Would any acidic or other undesirable products be formed that would adversely affect the 

liquid product quality and downstream processing? 

• Would any acid materials be formed which would cause corrosion of downstream 

recovery equipment? 

Co-injection of pyrolytic lignin in an FCC would be possible if the lignin was separated via water 

precipitation to keep the acid number below 35. The FCC is typically the largest unit in the 

refinery and so the lignin will be diluted to about 1.5% pyrolytic lignin in the VGO feed. For a 

50,000 bpd unit (typical for an average 150,000 bpd refinery) this would only allow 750 bpd of 

pyrolytic lignin in the feed. For a large 100,000 bpd FCC, 1500 bpd of pyrolytic lignin could be 

processed. Co-injection would also be possible if the pyrolytic lignin was previously hydrotreated 

to reduce the acid number. Since the FCC is the biggest unit and the heart of most refineries, 

much more development work would be required to minimize refiner risk before such an 

approach would be viable. 

30
 

IV.4 H2 PRODUCTION FROM PYROLYSIS OIL 

Pyrolytic lignin is deficient in hydrogen, making this 

portion of pyrolysis oil a poor feed for reforming 

into hydrogen. On the other hand, the organic 

compounds in the water-soluble phase are suitable 

for transformation into hydrogen. Table 20 shows 

the nominal composition used for the water-soluble 

phase organics on a moisture-free basis. 

Table 20. Composition of aqueousphase compounds from pyrolysis oil 

Component Wt% 

Glyoxal 10 

Formic acid 14 

Hydroxyacetaldehyde 17 

Acetic acid 17 

Ethylene Glycol 2 

Acetol 5 

Sugars and sugar anhydrides 35 

The following properties of the water-soluble phase 

of the pyrolysis oil must be considered in a H2-

production process: 

• The hydrogen content of the aqueous-phase 

organic compounds is substantially lower 

than that of typical hydrogen reforming 

feeds such as methane, LPG or naphtha. 

• The organic compounds for reforming are in a water solution. 

• The water-soluble phase consists of organic acids such as formic and acetic acids, with 

implications for processing unit metallurgy. 

• Higher molecular weight carbohydrates such as sugars and anhydro-sugars are likely to 

carbonize to tar or char while preheating to typical reforming temperatures. 

There are several technologies for converting hydrocarbons and biomasses into hydrogen. 

However, steam reforming is the primary technology for hydrogen generation in refineries and 

therefore the focus of the analysis of integrating pyrolysis oil into refining for hydrogen 

production (IV.4.1). The Aqueous Phase Reforming (APR) technology offered by Virent Energy 

Systems may be an alternate technology but most likely would be implemented in a stand-alone 

unit (Section IV.4.2). 

IV.4.1 Integration of Pyrolysis Oil Reforming into Refinery Hydrogen Plants 

The pyrolysis oil production basis for this study is 

7500 bpd of pyrolysis oil. The aqueous phase is 70 

lv% of the pyrolysis oil including water. The 

aqueous phase includes 30 lv% water. Therefore, 

approximately 3650bpd (60500lb/h) of aqueous 

phase components on a moisture free basis is 

available for reforming to hydrogen. 

Table 21. Pre-reformer syngas composition 

for the aqueous phase of pyrolysis oil. 

Adiabatic pre-reforming is the technology of choice 

to integrate reforming of the aqueous phase of 

pyrolysis oil into an existing hydrogen plant. For the 

nominal composition shown in Table 20 and 

assuming equilibrium at the reactor outlet, the 

resulting pre-reformer synthesis gas composition is 

found in Table 21. Figure 25 shows the proposed 

process flow scheme. The pre-reformer is an 

adiabatic fixed bed reactor operating at relatively 

low temperatures. Superheated steam is mixed with the preheated aqueous phase to produce a 

steam-oxygenates ratio of 1:1 (mass/mass). The preheated aqueous phase is mixed with the 

superheated steam as close as possible to the catalyst bed. Steam reforming, shift-gas and 

Component 

Composition 

(wt%) 

Sugars and Sugar 

Anhydrides 35.4 

Glyoxal 10.3 

Hydroxyacetaldehyde 16.9 

Formic Acid 14.1 

Acetic Acid 16.5 

Ethylene Glycol 2.3 

Acetol 4.5 

31
 

Aqueous 

Phase of 

Pyrolysis Oil

Convection

Superheated 

Steam

700 to 800°F

Boiler 

Feed

Water

Convection

Convection / 

Radiant

Steam Drum

Pre-Reformer 

SynGas

Figure 25. Adiabatic pre-reformer for the aqueous-phase of pyrolysis oil. 

methanation reactions occur in the adiabatic pre-reforming catalyst. Overall, the pre-reforming is 

highly exothermic for the aqueous phase composition. The pre-reformer synthesis gas is delivered 

without heat integration directly to the main steam reforming reactor. 

Based on the acidity of the feed, 317L stainless steel is the proposed metallurgy for the feed 

pump, fired heater, reactor cladding and associated piping and valves. The installed inside battery 

limits capital cost is approximately $15 to 20 MM, or $5500/bbl feed. 

Two cases were compared: 

• Case I - Hydrogen production from steam reforming of methane alone for the base 

refinery hydrogen production rate, 25MM scfd. 

• Case II - Reforming of the target quantity of the pyrolysis oil aqueous phase with 

adiabatic pre-reforming – 3650 bpd, 60,500 lb/h. Natural gas is supplemented in the main 

steam reforming plant as required to produce the baseline quantity of hydrogen. The 

steam to feed ratio is varied to observe the effect on H2 production and utility 

requirements. The effect on reforming catalyst stability is not addressed for the 

decreasing steam ratios. 

The analysis determined that the natural gas feed to the SMR reactor section can be completely 

displaced with the aqueous phase of the pyrolysis oil to produce at least the baseline H2

generation rate of 25 MMscfd. The CO2 emission per unit of hydrogen production is reduced by 

approximately 50 to 65% because components in the aqueous phase of the pyrolysis oil do not 

contribute to CO2 emissions while all carbon in the natural gas feed for SMR is eventually 

emitted as CO2. 

32
 

The hydrogen COP is approximately 50% higher for integrated adiabatic reforming of the 

aqueous phase of pyrolysis oil than for SMR. Feed costs represent two-thirds of the hydrogen 

COP and almost the entire COP for SMR. The projected decrease in CO2 emission per unit of 

hydrogen is not sufficient to make up the difference through CO2 credits. The delivered cost of 

pyrolysis oil must be reduced for this process to be competitive with SMR. 

IV.4.2 Aqueous Phase Reforming (APR) of the Aqueous Phase of Pyrolysis Oil 

The Aqueous Phase Reforming (APR) process is a concept patented by Virent Energy Systems. A 

parametric simulation study was performed examining the effects of reaction temperature and 

pressure on hydrogen generation and cost of production for APR of pyrolysis oil with the 

composition shown in Table 20. The ranges of temperature and pressure of interest were those 

cited by the patent and other public literature from Virent Energy Systems: 350 to 500°F and less 

than 1000 psia. APR was modeled with the PRSV thermodynamic property package in a process 

simulator using reforming equilibria for the reactions in Table 22. The effects of temperature and 

pressure on hydrogen generation indicate that the extent of reaction is favored by increasing 

temperature and pressure. Vaporization of the feed in the fired heater is energy intensive. 

Therefore, in this analysis, the upper limit of temperature was bounded by the saturated liquid 

temperature at a given pressure. 

A proposed flow scheme for the APR process is shown in Figure 26. The aqueous phase of the 

pyrolysis oil and recycled water are heated up to reaction conditions through exchange with the 

reactor effluent and a fired heater. Pressure, temperature and the water recycle rate can be 

optimized to economically maximize the conversion to hydrogen. After exchanging heat with the 

reactor feed, the reactor effluent is cooled by an air cooler and the predominantly liquid recycle is 

separated from the synthesis gas. The synthesis gas is primarily composed of carbon oxides and 

hydrogen. A Pressure Swing Adsorption (PSA) unit is used to purify the H2 and reject a tail gas. 

Table 22. Reforming equilibrium reactions for 

APR of pyrolysis oil components 

Component Reforming Equilibrium Reaction 

Acetic Acid Reforming CH3(COOH) = 2 H2 + 2 CO 

Acetol Reforming CH3(CO)CH2OH = 4 H2 + 3 CO 

Ethylene Glycol Reforming H2C(OH)C(OH)H2 = 3 H2 + 2 CO 

Formic Acid Reforming H(COOH) = H2O + CO 

Glyoxal Reforming H(CO)(CO)H = H2 + 2 CO 

Hydroxyacetaldehyde Reforming H2C(OH)(CO)H = 2 H2 + 2 CO 

Sugar Reforming C6H12O6 = 6 H2 + 6 CO 

Water Gas Shift CO + H2O = H2 + CO2

Process modeling suggests that part of the plant will be exposed to hot acetic vapor which poses 

metallurgical challenges. The corrosivity of acetic acid solutions (and other organic acids like 

formic acid) is very dependent on process conditions and existing contaminants. 316 SS may not 

be an acceptable metal of design for boiling acetic acid solutions at concentrations of 10% or 

higher. Copper or alloys of copper and nickel may be viable alternatives, especially in oxygenfree environments, unless the material is transferred from tankage where oxygen content is a real 

concern. Development of a commercial APR process will require corrosion testing at the various 

process conditions found throughout the process. 

33
 

Aqueous 

Phase of 

Pyrolysis Oil

PSA

H2

Reactor

Separator

Tail 

Gas

Unconverted 

Pyrolysis Oil Water Recycle

Figure 26. Proposed APR process flow scheme. 

The most conservative approach is to use titanium Grade-II clad carbon steel. The difference in 

erected cost between Ti-clad and 316SS is surprisingly small - within 5%. The total installed 

inside battery limits capital cost is approximately $65 MM for the APR plant based on the 

proposed flow scheme and the baseline hydrogen production rate of 25 MMscfd. 

The hydrogen COP from pyrolysis oil aqueous 

phase reforming is appreciably higher than 

SMR. Alternate design concepts were 

considered to increase economic feasibility by 

reducing capital investment and COP. The most 

promising is similar to the flow scheme of 

Figure 26, but without liquid recycle. A 

summary of the COP and estimated capital costs 

comparing the initially proposed flow scheme 

and the low acetic-acid conversion alternative 

are presented in Table 23, assuming $16/bbl for 

pyrolysis oil. The process results in low acetic 

acid conversion and includes recovery of the 

acetic acid as a net product. Two distillation 

columns could be used for acetic acid recovery. Waste water and non-condensable light ends are 

vented overhead in the first column. Purified acetic acid is recovered in the overhead of the 

second column; tar or other heavier by-products are removed in the bottoms. With elimination of 

recycle and assuming efficient oxygenate reforming of the feed pyrolysis oil, the COP can be 

reduced to approximately $0.48/lb, which includes acetic acid fractionation utilities. Acetic acid 

fractionation costs are approximately $0.06/lb H2 generated. 

Table 23. Comparison of 

APR process flow schemes. 

Process Concept COP 

($/lb H2) 

EEC 

(MM $US) 

Preliminary 

Process Design – 

complete acetic 

acid conversion 

process 

0.65 65 

Low acetic acid 

conversion process 

0.48 10 

34
 

Table 24 compares the capital costs and 

production capacity for the three routes 

to generate hydrogen from the watersoluble portion of pyrolysis oil: steam 

methane reforming, partial oxidation 

and liquid phase reforming. All are 

based on 7,500 bpd pyrolysis oil feed 

that has been separated into the lignin 

and water-soluble fractions. 

Table 24. Comparison of H2 production routes. 

EEC Capital 

Cost ($MM) 

Production Capacity 

(MMSCFD) 

Steam reforming 24 28.7 

Partial oxidation 37 15.9 

Liquid phase 

reforming 30 33.3 

IV.5 HYDROTHERMAL PROCESSING AND BLENDING STUDIES

Supporting studies were conducted to investigate the potential for pretreating pyrolysis oil by 

hydrothermal processing and for improvements to pyrolysis oil blending properties. 

Hydrothermal Processing 

Hydrothermal processing is a simple pretreatment step which could be done in the field to reduce 

the costs of transporting pyrolysis oil by removing oxygen and reducing acidity. Batch autoclave 

tests were performed to hydrothermally process the pyrolytic lignin, pyrolysis oil and resid. In 

these tests, water is mixed with pyrolytic lignin and the mixture is heated up to 300 or 350 C. The 

pressure is maintained so that the mixture is kept in the liquid phase. The experimental conditions 

are shown in Table 25. The catalysts were cobalt and nickel compounds. 

Table 25. Hydrothermal processing experiment conditions. 

Pyrolytic lignin,g Resid g Catalyst Water,g Temp C Time min Pres Psi (max) Water phase,g Organic phase Gas,l Gas,% CO2 

1 75.7 0 75.1 300 15 1380 69.9 69.9 2.2 33 

2 51.5 50.8 50.1 350 15 1530 5 131.2 2.2 12 

3 76.6 75.2 350 15 2500 20.5 107.2 11 

4 52.5 60.7 50.6 350 120 2480 28.6 106.6 2.9 24 

5 75.7 75.0 350 120 2480 29.3 105.1 .03 6 

6 77.8 Co >76.7 350 20 2850 186 

total 2.8 55 

7 77.0 Ni 75.1 350 20 2800 77.0 46.5 3.3 47 

8 75.1 75.3 300 15 2100? 74.3 59.5 2.1 27 

9 75.0 76 

(pitch) 250 15 480? 145 1.1 45 

35
 

The resid did not appear to improve with hydrothermal treatment, but the pyrolytic lignin lost 

oxygen as shown in Table 26. Unfortunately the product from hydrothermal treatment became 

more viscous as well (>1500 MPas @ 95 C) and some char formation is reported in Runs 6,7and 

8. Hydrotreatment under H2 pressure breaks the molecules as the oxygen groups are removed, 

resulting in smaller molecules. Hydrothermal treatment does not have this effect and may result 

in polymerization of the molecules. As long as the aromatic rings have not condensed to form 

asphaltene structures, it may be possible to hydrocrack these materials to form aromatics in a 

hydrocracker. A key test would be to determine whether hydrothermally-treated pyrolytic lignin 

can be processed with LCO in a mild hydrocracker to produce aromatics. Since the hydrothermal 

process is a simple and relatively inexpensive option for pretreatment, more follow up work in 

this area is recommended. 

Blending Studies 

One of the big challenges with 

pyrolysis oils is the fact that they have 

limited solubility in typical 

hydrocarbon fractions and fuel 

products. Small amounts of materials 

can be extracted from pyrolysis oils 

into hydrocarbon phases but blending 

studies are difficult because the 

pyrolytic lignin or other heavy 

fractions of pyrolysis oils tend to coat 

onto beakers and equipment leading to 

poor material balance. Therefore to be 

conservative, only increases in the 

petroleum fraction phase (VGO or 

LCO) are counted as extracted 

material. Roughly 3% pyrolytic lignin was extracted into LCO (light cycle oil) in blending tests 

and roughly 6% pyrolysis oil could be extracted into VGO when water was added to the pyrolysis 

oil phase. 

Table 26. Product properties of hydrothermally 

processed pyrolytic lignin. 

Pyrolytic 

lignin 

Hydrothermally 

treated lignin 

Process conditions 

Co-feed 50% water 

Temp,C 300 

Time 15 min 

%C 56.67 79.1 

%H 6.67 13.5 

%O 26.6 7.1 

%N 0.2 

%water 11.7 

However it was found that stable microemulsions of 20% pyrolysis oil or 20% pryolytic lignin in 

VGO or LCO could be formed when 1% Hypermer 2296 and 1% Hypermer 256SF surfactants 

were used. Also encouraging was the fact that the HT-82 sample of hydrotreated pyrolytic lignin 

was very soluble in both VGO and LCO. 

IV.6 ECONOMIC ANALYSIS OF PYROLYSIS OIL PROCESSING OPTIONS

Table 27 shows that pyrolysis oil is 

intermediate in cost between heavy residual 

oil and natural gas on a $/MMBTU basis. 

Refiners could burn pyrolysis oil in furnaces 

designed for heavy fuel oil that have been 

shifted to natural gas to minimize SO2

emissions. This would also allow refiners to 

reduce CO2 emissions from boilers and fired 

heaters and obtain a CO2 credit. This is an immediate application for refiners and other industries 

using natural gas to minimize pollution. Minor capital investment would probably be needed to 

change burner tips and modify boilers to allow co-firing of natural gas and pyrolysis oil. It would 

be worth $5.6 MM/year to substitute 5250 bbl/d of water soluble pyrolysis oil for natural gas in 

furnaces or boilers or $7.4 MM/year to substitute 5250 bbl/d of pyrolysis oil itself. 

Table 27. Fuel value of pyrolysis oil. 

$/MMBTU 

Heavy Fuel Oil 4.2 

Natural Gas 6 

Pyrolysis Oil 4.7 

36
 

The baseline 5250 bpd of pyrolysis oil aqueous phase contains 3650 bpd water-soluble 

components on moisture free (m.f.) basis. Assuming the nominal composition in Table 20, the 

mass rate is 65,940 lb/h and the lower heating value of the m.f. organics is approximately 5600 

Btu/lb. The water in this stream has no net heat value. 

To raise 1600 psig, 700°F VHP superheated steam from ambient temperature boiler feed water 

requires 1.3 MMBtu per 1000lb steam. Therefore, burning the aqueous phase generates 

370MMBtu/h on a 100% efficiency basis. The efficiency basis is valid if the aqueous phase of the 

pyrolysis oil provides supplemental firing for the heat recovery-steam generation (HRSG) section 

of a cogeneration facility. Otherwise, 85 to 90% efficiency in a stand-alone boiler would be more 

appropriate. The amount of VHP superheated steam generated will be approximately 250,000 to 

285,000 lb/h. The steam generated is adequate for power generation or the steam could be 

generated at HP steam levels to drive 1 or 2 large refinery condensing turbines. 

The economics for producing gasoline from pyrolytic lignin are shown in Table 28, assuming 

$18/bbl pyrolysis oil ($16/bbl +$2/bbl transportation charges) and $40/bbl crude. The incentives 

total to 4 MM/year. 

Table 28. Incentives for gasoline 

production from pyrolytic lignin. 

$/D Quantities 

Feed 

Pyrolytic lignin 40,500 2250 bpd 

H2 25,680 21.4 tpd 

Product 

Light ends 19,903d 64 tpd 

Gasoline 52,520 1,010 bpd 

Diesel 12,000 250 bpd 

Utilities -4,800 

Incentive 12,843 

The capital cost for a 

hydrotreating/hydrocracking unit 

producing 1,010 bpd of gasoline is 

$30 MM. Producing gasoline from 

pyrolytic lignin becomes 

economically attractive for $18/bbl 

pyrolysis oil at a crude price of 

$50/bbl. As pyrolysis oil price 

decreases, producing gasoline from 

pyrolytic lignin becomes attractive at 

lower crude oil prices, illustrated in 

Figure 27 .There is a strong sensitivity 

to pyrolysis oil price at $40/bbl crude 

but the sensitivity decreases as crude 

oil price increases. 

0

1

2

3

4

5

6

7

8

$40/bbl $50/bbl $60/bbl

$18/bbl Pyrolysis oil

$14/bbl Pyrolysis oil

Years

0

1

2

3

4

5

6

7

8

$40/bbl $50/bbl $60/bbl

$18/bbl Pyrolysis oil

$14/bbl Pyrolysis oil

Years

Figure 27. Years to payback for conversion of pyrolytic 

lignin to gasoline. 

37
 

Table 29 shows the incentives for producing H2

from the water-soluble portion of pyrolysis oil for 

$40/bbl crude and $10/bbl pyrolysis oil. Pyrolysis 

oil must be less than $12/bbl before there is a 

positive incentive for H2 production. The incentive 

is $3.5MM/year with the CO2 credit and 

$0.3MM/year without. 

Table 29. Incentives for H2 production 

$/D Quantities 

Feed 

Water-soluble pyrolysis oil 52,500 5.250 bpd 

Products 

Hydrogen 84,910 28 MMSCFD 

CO2 credit ($30/ton) 9,770 

Utilities -31,500 

Incentive 10,680 

The capital cost of a pre-reformer to add to an 

existing H2 plant is $15MM. Unless pyrolysis oil 

price drops, it is better for a refiner to simply co-fire 

the aqueous phase pyrolysis oil as fuel, replacing 

natural gas which is used in the refinery. 

Table 30. Overall incentives for use of pyrolysis oil for 

fuel and gasoline production. 

$/D Quantity

Feed 

Pyrolysis oil 135,000 7500 bpd

Water Soluble Pyrolysis Oil 5250 bpd

Pyrolytic Lignin 2250 bpd

Products

Fuel 66174 5250 bpd

Lt Hydrocarbons 19,303 64 tpd

Gasoline 52,520 1010 bpd

Diesel 12,000 250 bpd

CO2 credit ($30/ton) 19,060

Utilities - 4800

Incentive 29257

Table 30 shows the economics of cofiring 

the water soluble portion of pyrolysis oil 

and making gasoline from the pyrolyic 

lignin. Figure 28 shows the time to 

payback for conversion of the pyrolytic 

lignin to gasoline and co-firing of the 

water soluble pryolysis oil as a fuel 

substitute for natural gas. 

 

0

1

2

3

4

5

6

7

8

9

10

$40/bbl $50/bbl $60/bbl

CO2 credit

No CO2 Credits

Years

0

1

2

3

4

5

6

7

8

9

10

$40/bbl $50/bbl $60/bbl

CO2 credit

No CO2 Credits

Years

Figure 28. Years to payback for combined process. 

38
 

V OPPORTUNITIES FOR FISCHER TROPSCH LIQUIDS

Another option for biorenewable use is gasification followed by Fischer-Tropsch (FT) synthesis 

from the syngas. The properties of crude FT liquids are summarized in Table 31 and Table 32. 

Crude FT liquids contain some oxygen and olefins which are easily removed by hydrotreating. 

The study calculated the value of FT liquid to a typical refiner if biorefineries chose to send them 

to the refinery rather than upgrade the FT themselves. The yields and properties are based on an 

iron fixed bed FT system marketed by Rentech. 

Table 31 Properties of crude FT liquids Table 32. Detailed properties of crude FT liquids. 

Crude Resid Crude FT 

Liquids 

% C 83-86 84.9 82.8 

%H 11-14 10.6 14.0 

%S 0-4 

(1.8avg) 

4.2 - 

%N 0-1 

(.1avg) 

.3 - 

%O - - 3.2 

H/C 1.8-1.9 1.5 2.0 

Density .86(avg) 1.05 0.82 

TAN # <1 <1 Low 

Heating 

value 

BTU/lb 

18,000 17,500 18,400 

Naphtha Diesel Wax 

Wt% 19 21 60 

Density 0.76 0.79 0.85 

BP, F 70 – 320 320 – 650 650+ 

Wt% Oxygen 6 6 1.3 

Wt% Saturates 22 30 88 

Wt% Olefins 48 14 6 

Wt% Alcohols 26 16 6 

The configuration of a typical hydrocracking-based refinery was shown in Figure 5 of Section 

II.2.3. In order to maximize diesel production, the wax is cracked in a hydrocracker (HCK) to 

make diesel and the naphtha and diesel must be hydrotreated to remove oxygen and olefins. 

There are a number of options possible for integration with a typical refinery. These include: 

• Send FT Liquids to the crude tower if all diesel is hydrotreated. This places a penalty on 

gasoline produced (poor octane) but requires no capital. 

• Send everything to the HCK. This results in yield loss from cracking diesel but requires 

no capital. 

• Send FT liquids to their own tower, while sending the diesel and naphtha cut to DHT and 

the wax cut to HCK to produce diesel and naphtha. Capital is required for two towers and 

typical HCK conditions are not optimal for FT wax. 

• Send FT liquids to their own tower, while sending the diesel and naphtha cut to DHT and 

the wax cut to low pressure HCK (converted DHT). Capital is required for two towers if 

DHT is available for use as HCK (optimal configuration). 

Sending FT their own tower and processing the wax in the hydrocracker and the naphtha and 

diesel in the DHT is the best option, but does require a small investment in two distillation 

39
 

towers. The resulting configuration is shown in Figure 29. The value of the FT liquids to the 

refiner is shown in Table 33. 

HCK

SGP SGP

FCCU

VISB

ISOM

LPG

Butanes

Gasoline

Jet

CSO

Diesel

CH4

Fuel Oil

LCO

NHT

MEROX

C

D

U

Crude

Oil

All H2

Users

MTBE

ALKY

F

F

F

H2 Plant

V

D

U

FT

Liquid

Diesel 

Hydrotreating

Naphtha

Reformer

HCK

SGP SGP

FCCU

VISB

ISOM

LPG

Butanes

Gasoline

Jet

CSO

Diesel

CH4

Fuel Oil

LCO

NHT

MEROX

C

D

U

Crude

Oil

All H2

Users

MTBE

ALKY

F

F

F

H2 Plant

V

D

U

FT

Liquid

Diesel 

Hydrotreating

Naphtha

Reformer

Figure 29. Typical hydrocracking refinery processing crude FT liquids. 

Table 33. Value of crude FT liquids to refiners. 

Starting 

wt%

Wt% after 

conversion

Value 

$/ton

Total 

Value 

$/ton

Lt ends 3.7 320 11.8 

Naphtha 19 28.9 324 93.6 

Diesel 21 66.4 368 244.4 

Wax 60 

Water 3.5 

H2 use -1.4 1200 -16.4 

Operating cost and 

refinery margin -20 

Total value $/ton 313 

40
 

VI ACRONYMS AND ABBREVIATIONS. 

Table 34 lists the acronyms and abbreviations used elsewhere in this report. 

Table 34. Acronyms and abbreviations. 

Acronym or 

Abbreviation 

Definition 

ABT Average bed temperature 

APR Aqueous phase reforming 

bbl Barrel(s) 

bpd Barrels per day 

bpsd Barrels per stream day 

Br No Bromine Number 

COP Cost of production 

CSO Clarified Slurry Oil 

DeCO2 Decarboxylation 

DHT Distillate hydrotreater 

EEC Estimated erected cost 

EOR End of run 

FAME Fatty acid methyl esters 

FCC Fluid catalytic crackers 

FFA Free fatty acid(s) 

FT Fischer-Tropsch 

GJ Gigajoule = 1.055 MMBtu 

HC Hydrocarbon(s) 

HCK Hydrocracking 

HDM Hydrodemetallization 

HDO hydrodeoxygenation 

HDS Hydrodesulfurization 

HP High pressure 

HT Hydrothermal 

HT Hydrotreating 

LCA Life-cycle analysis 

LCO Light cycle oil 

LHSV Liquid hourly space velocity 

LHV Lower heating value 

LPG Liquified petroleum gas 

lv% Liquid volume % 

MCFD Million cubic feet per day 

mf, m.f. Moisture-free 

41
 

MMBtu Million Btu 

MTBE Methyl tertiary butyl ether 

NPV Net present value 

POP Proof of principle 

POX Partial oxidation 

PRSV Peng-Robinson-Stryjek-Vara 

PSA Pressure swing adsorption 

scfb Standard cubic feet per barrel 

SMR Steam methane reforming 

SOR Start of run 

SS Stainless steel 

TAN Total acid number 

tpd Tons per day 

ULSD Ultra-low sulfur diesel 

VGO Vacuum gas oil 

VHP Very high pressure 

WHSV Weight hourly space velocity 

WTI West Texas Intermediate 

42
 

VII REFERENCES

 1

 DOE/USDA. “Biomass as a Feedstock for a Bioenergy and BioProducts Industry: The Technical 

Feasibility of a Billion-Ton Annual Supply”, April, 2005, pp xii and 3. Joint study by DOE and USDA. 

Available at: http://www.eere.energy.gov/biomass/pdfs/final_billionton_vision_report2.pdf. 2

 Growing Energy: How Biofuels Can Help End America’s Oil Dependence, pg 35. Published by NRDC, 

December 2004. Available at http://www.bioproducts-bioenergy.gov/pdfs/NRDC-Growing-EnergyFinal.3.pdf. 3

 Lynd, et al. in World Congress on Industrial Biotech and Bioprocessing, Orlando, FL, April 20-22, 2005. 

(Liquid Transportation Fuels). 

4

 Tyson, K. Shaine, “Oil and Fat R&D”. Presentation by NREL to UOP, July 17, 2003. 

5

 Tyson, K. Shaine, Bozeli,J, Wallace,R., Petersen Eugene, Moens,Lue, “Biomass Oil: Analysis Research 

Needs and Recommendations” NREL/TP-510-34796, June 2004. 

6

 Pervin and Gertz, Inc. (Subscription only). www.pervingertz.com. 7

 Larsen, Eric D, et al.. Energy for Sustainable Development, 2000. “Expanding roles for modernized 

biomass energy” Energy for Sustainable Development, Volume IV, No. 3, October 2000. 

8

 DOE/USDA. op. cit., pg 105. 

9 http://www2.barchart.com/comfund/tallow.asp. 10 Spath, P. L., Dayton, D. C. (2003) Preliminary Screening – Technical and Economic assessment of 

synthesis Gas to Fuels and Chemicals with Emphasis on the Potential for Biomass-Derived Syngas. 

NREL/TP-510-34929. 

11 Grange,P, Laureat,E.,Maggi,R, Centeno,A.,Delmon,B, “Hydrotreatment of pyrolysis oil from 

biomass:reactivity of the various categories of oxgenated compounds and preliminary technico-economic 

study , Catalysis Today 29 (1996)297-301. 

12 Maier, et al. Chem. Ber. 115, p. 808 – 812 (1982). 

13 Arbokem, U.S. patent #4,992,605 

14 Canmet/Arbokem U.S. Patent #4,992,605 

15 Radich, Anthony. Biodiesel Performance, Costs, and Use. Available at 

http://www.eia.doe.gov/oiaf/analysispaper/biodiesel/. 

16 USDA. USDA Agricultural Baseline Projections to 2012, Staff Report WAOB-2003-1 (Washington, 

D.C. February 2003). Available at www.ers.usda.gov/publications/waob031/waob20031.pdf. Cited by 

Energy Information Administration , Office of Integrated Analysis and Forecasting. 

17 National Biodiesel Board. Tax Incentive Fact Sheet. Available to members at: 

https://www.biodiesel.org/members/membersonly/files/pdf/fedreg/20041022_Tax_Incentive_Fact_Sheet.p

df 

18 Aden, A. Memorandum prepared for UOP by NREL: “Biodiesel Information for UOP” January 18 – 19, 

2005. 

19 www.uniqema.com/og/lit/hypermersurfactants.pdf. 20Samolada MC and Vasolos IA. Catalytic Cracking of Biomass Flash Pyrolysis Liquids” in Developments 

in Thermochemical Biomass Conversion, in Proceedings of the International Conference on Developments 

in Thermochemical Biomass Conversion. 

21 Samolada MC and Vasolos IA. Production of Transportation Fuels by Upgrading of Biomass Flash 

Pyrolysis Liquids via FCC Technology” in Proceedings of the 9th European Bioenergy Conference 24-27 

June 1996. Pergamon 1997, pp 1578-1583. 

22 Bormann, Konrad; Tilgner, Hartmut. Influence of rape seed oil triglycerides on the catalytic cracking of 

petroleum distillate. Erdoel, Erdgas, Kohle (1994), 110(2), 75-7. 

43
OPPORTUNITIES FOR 

BIORENEWABLES IN OIL REFINERIES

APPENDIX:

BACKGROUND AND EXPERIMENT SUMMARY

PREPARED FOR DOE 

uop

25 East Algonquin Road 

Des Plaines, Illinois 60017-5017 

U.S.A. 
 

A select set of experiments performed in the project are summarized in Table 1 and Table 2. 

Table 1. Experiment summary (Part I). 

524 529 531 535 537 538 540 

Feed 

vegetable oil 

HT-82 

( HT lignin) 

Pyrolytic lignin 

(Dynamotive) Pyrolysis Oil

Conditions 

Temperature 300-350 300-350 300-350 370 350 350 350 

Time at temp hr 2.1 1.88 1.7 2 3 4 3 

Catalyst AS-250 HCT-1 N200 HC43 HCT HCT HCT 

Cat amt ,g 25 25 25 25 25 25 25 

Feed amt 75 75 75 75 75 150 75 

WHSV 1.4 1.6 1.76 1.5 1 1.5 1 

LHSV 0.91 1.40 1.26 0.96 0.67 1.01 0.67 

Pressure psig 500 500 500 1500 1500 1500 1500 

Products 

Water % of feed 2 1.7 1.15 7.5 24.1 23.8 33 

Water from organic % of 

feed 

4.6 12.4 12.1 6.3 

Light liquids % 4.8 0.09 0.64 30 27.7 19.8 19 

Diesel % 65.07 78.5 76.4 29.4 12.8 28.8 19.9 

VGO % (necessary to remove to 

reach diesel spec) 

3.2 

Total % liquid product 69.87 78.59 80.24 59.4 40.5 48.6 38.9 

Gas and lights( by 

difference) 

28.13 19.71 18.61 33.1 35.4 27.6 28.1 

Calculated CO2+CO 10.9 12.7 13.4 18.4 24.0 18.6 37.1 

Lt Hc and loss calc 17.3 7.0 5.2 14.7 11.4 9.0 -9.0 

Total 100 100 100 100 100 100 100 

% triglyceride conversion 100 100 100 

% O in liquid product 3.4 1.3 0.9 1.8 4.8 7.3 5.5 

%O in naphtha 1.3 5.9 7.2 8.6 

%Oxygen in diesel+ 3.4 1.3 0.9 2.3 2.4 7.3 2.4 

%Oxygen removal 81.2 90.1 91.3 94.5 94 90 96 

% O to water 15 12.8 9.8 34.36 58.53 57.80 53.33 

% Organic O to water 31.79 48.13 47.40 29.60 

% oxgen to Co and Co2 66.2 77.3 81.5 60.14 41.47 32.20 42.67 

% feed to CO2(calc) 10.9 12.7 13.4 18.4 24.0 18.6 37.1 

%organic Oxygen to Co+Co2 62.7 45.9 42.6 66.4 

%NC15 6 9 8.9 

%IC15 2 1.2 

%NC16 3 3 4.6 

%IC16 2 0.5 1.4 

%NC17 37 55 43.6 

%iC17 1.8 2 4.1 

%NC18 19.6 18.5 21.1 

%iC18 2 3.5 5.33 

Total 73.4 91.5 90.23 

Feed %Oxygen 10.4 10.4 10.4 19.4 36.6 36.6 57 

Feed density 0.92 0.92 0.92 1.03 1.2 1.2 1.2 

Acid number feed 34 168 168 75 

Water in feed 2.9 11.7 11.7 26.7 

Feed IBP 21 

Feed 10% 126.6 

Feed 30% 225.4 

Feed 50% 323 

Feed 70% 413 

Feed 90% 529 

Feed EP 607 

2
 

Table 2. Experiment summary (Part II). 

524 529 531 535 537 538 540 

Feed vegetable oil 

HT-82 

( HT lignin)

Pyrolytic lignin 

(Dynamotive) Pyrolysis Oil 

Acid number product 

Acid number trap HC 15.1 24.11 43 

Acid number diesel 0 

Naphtha density 0.81 

Diesel density 

Distillations 

 Product IBP C -8 

Product wt 10% 105 

Product 30% 157 

Product 50% 199.5 

Product 70% 246.5 

Product 90% 332 

Product EP 592 

%m TOC in water 2.4 4.1 3.2 

naphtha fischer water 946ppm 

Naphtha IBPC -8 20 -10 25 

10% 71 80 22 86 

30% 105 133 105 147 

50% 143 173 109 195 

70% 181 216 138 243 

90% 226 277 201 310 

EP 321 459 312 422 

Naptha %C 86.1 82.7 78.4 78.6 

Naptha %H 12.6 11.4 10.1 11 

Naphtha %O by dif 1.3 5.9 11.5 10.4 

Diesel_ IBP 50 24 

10% 139 267 

30% 209 319 

50% 256 358 

70% 312 413 

90% 438 500 

EP 592 600 

Diesel %C 82.8 84.5 84.7 87.1 87.2 83.9 87.6 

Diesel %H 13.7 14.2 14.4 10.3 10.4 8.8 10 

Diesel %O(by dif) 3.5 1.3 0.9 2.6 1 7.3 2.4 

Comment - Lots of 

cracking 

- No gas 

samples 

- Some 

isom. 

- Less 

cracking 

- Some 

gas 

samples 

- Less 

isom. 

Too much 

heavy 

ends 

some 

coking 

3
 

We provide here supporting information for the study basis outlined in Section II.2 of the report. 

Table 3 lists some of the significant U.S. and international mandates that will drive increased use 

of biofuels in the next few years. 

Table 3. Global mandates for biofuel usage. 

2005 2010 Type 2010 Targets 

(bbl/d) 

EU 2% biofuels 5.8% biofuels Law 200,000 

U.S. diesel 

No % but high 

subsidies for 

biodiesel 

2.5% Goal 100,000 

U.S. ethanol 2.2% of gasoline 3.5% of 

gasoline Goal + Law 325,000 

India 5-10% Goal 80,000-160,000 

4
 

Table 4 lists the base prices of feeds used in the study. Bio-feedstock prices were based upon an 

average of 2003 – 2004 values reported in the Chemical Marketing reporter. 

Table 4. Base prices.

$/unit $/ton 

WTI Crude 40/bbl 306 

Arab Medium 36.5/bbl 237 

Reg Gasoline 49/bbl 410 

Premium gasoline 52/bbl 435 

Diesel 48/bbl 343 

Residual Oil 27/bbl 173 

Methanol 34/bbl 275 

Soy Oil without subsidy 0.25/lb ($73/bbl) 559 

Soy Oil with subsidy 0.06/lb ($17/bbl) 132 

Yellow Grease without subsidy 0.14/lb ($41/bbl) 310 

Yellow grease with subsidy 0.02/lb ($6/bbl) 44 

Brown Grease 0.06/lb ($18/bbl) 132 

Pyrolysis Oil 16/bbl 83 

Glycerol 0.10/lb 220 

Methane( Natural Gas) 6.0/MMBTU 286 

Toluene 62/bbl 450 

Ethylene 695 

Propylene 670 

Propane 31/bbl 352 

H2 3.1/MSCF 1200 

Benzene 700 

Xylene 575 

Crude tall oil 100 

5
 

Subsidies were calculated based on the subsidies for biodiesel shown in Table 5.

1

 

Table 5. Biodiesel subsidies (feedstock prices from crop year 2003/04 projections). 

Feedstock 

Feedstock 

Price,$/lb 

Total 

Production 

Cost, 

$/gal B100 

no incentives

EstimatedCCC 

Incentive 

Net Biodiesel 

Production 

Cost, 

$/gal B100 

Net Biodiesel 

$/gal B20 

with tax credit

Soy 0.238 2.29 1.43 0.86 0.66 

Sunflower 0.262 2.49 1.57 0.91 0.71 

Corn 0.263 2.49 1.58 0.92 0.72 

Peanut 0.436 3.9 2.62 1.28 1.08 

Cottonseed 0.279 2.62 1.67 0.95 0.75 

Safflower 0.735 6.32 4.41 1.91 1.71 

Edible Tallow 0.164 1.69 0.98 0.71 0.61 

Inedible 

tallow 

0.151 1.59 0.91 0.68 0.58 

Lard 0.179 1.81 1.07 0.74 0.64 

Choice white 

grease 

0.129 1.41 0.77 0.64 0.54 

Yellow grease 0.104 1.21 0.62 0.58 0.48 

 1

 NREL. Presentation on Biodiesel Costs. July, 2003. 

6
 

Table 6 and Table 7 list the costs of pyrolysis oil and related assumptions. 

Table 6. Pyrolysis oil costs. 

Biomass 

Federate 

Biomass 

Feed Cost Pyrolysis Oil Produced Pyrolysis Oil Cost 

Tonne/day $/dry ton t/day bbl/day $/GJ (LHV) $/MMBtu $/BBL $/BBL of FOE 

500 $30.00 325 1705 $6.77 $6.42 $21.10 $51.79 

500 $46.00 325 1705 $8.20 $7.77 $25.53 $62.67 

2000 $30.00 1300 6820 $5.10 $4.84 $15.89 $39.01 

2000 $46.00 1300 6820 $6.52 $6.18 $20.32 $49.88 

Table 7. Additional assumptions. 

1 GJ 1.055 MMBtu 

1 Barrel 42 gallons 

Density of Pyrolysis Oil 4.55 kg/gal 

LHV of Pyrolysis Oil 16,300 kJ/kg 

LHV of fuel oil 

(No. 6 Fuel Oil from Phyllis) 40,009 kJ/kg 

7
 

The properties of typical biofeedstocks are summarized in Table 8. The low heating value for 

pyrolysis oil is the result of the high oxygen content. The high alkali metals in vegetable oil, 

grease and pyrolysis oil are mostly potassium and calcium. 

Table 8- Feedstock properties 

Crude Resid Pyrolysis Oil Soy Oil Yellow 

Grease 

% C 83-86 84.9 56.2 77/6 76.4 

%H 11-14 10.6 6.6 11.7 11.6 

%S 0-4 (1.8avg) 4.2 - 0.0006 0.04 

%N 0-1 (0.1avg) .3 0.3 0.0011 0.03 

%O - - 36.9 10.4 12.1 

H/C 1.8-1.9 1.5 1.4 1.8 1.8 

Density .86(avg) 1.05 1.23 0.92 0.89 

TAN # <1 <1 78 2 30 

Ppm alkali 

metals 60 6 100 100 100 

Heating value 

BTU/lb 18,000 17,500 6,560 16,000 16,000 

8
 

The relative locations of Biofeedstock sources and refineries in the U.S are shown in Figures 1 

and 2. 

Oil and Grease production

Refineries

Figure 1. Locations of U.S oil, grease and refineries.

Agricultural resources

Wood resources 

Agricultural and 

Low inventory

residues

and residues

wood residues

refinery

Figure 2. Locations of U.S. biomass and refineries. 

9