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# O PPORTUNITIES FOR 

# B IORENEWABLES IN 

# O IL 

# R EFINERIES



# F INAL 

# T ECHNICAL 

# R EPORT

S UBMITTED TO 

# : U.S. D EPARTMENT OF 

# E NERGY



# 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



# T

ABLE OF 

# C ONTENTS

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 S upplies ...................................................................................................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 Me thyl Ester) Production...............................................................8 III.1.1 Fatty Acid Methyl Ester Production in Existing Refi ning Units ...............................9 III.1.2 Retrofit E nablers...................................................................................................11 III.1.3 Conclusions ..........................................................................................................13 III.2 Green Diesel Production...............................................................................................14 III.2.1 Green Diesel Product ion Approaches..................................................................14 III.2.2 Green Diesel Production – Co-Processing Biof eedstocks in an Existing Unit.....14 III.2.3 Green Diesel Production – Processing Biofeedstocks in a Stand-Alone Unit......18 III.2.4 Green Diesel Experim ental Resu lts .....................................................................20 III.3 Comparison of Biodiesel and Green Diesel Process and Properties ...........................21 III.4 Green Gasoline and Green Olefins Produ ction ............................................................22 III.5 Economic Assessment and Li fe Cycle A nalysis ...........................................................23 

### IV Refinery Opportunities for Pyrolysis Oil ............................................................ 25
IV.1 Background and Proces sing Opti ons............................................................................25 IV.2 Gasoline And Aromatics Producti on from Pyroly tic 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 Reformi ng into Refinery Hydrogen Pl ants.................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 Pyrolysi s 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



# L IST OF 

# F IGURES 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



# L IST OF 

# T ABLES 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 hydrotr eating/hydrocracking pyrolytic lignin. ..................... 28 Table 17. Yields for FCC conversion of pyrolysi s 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 H 2 

### 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 H 2 

### 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 E XECUTIVE 

# S UMMARY
The purpose of this study was to evaluate the poten tial 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 pr oof 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 utili zation 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 pro duce, 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 vegetabl e 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, particul arly if eligible for subsidies. Processing of pyrolysis oils represents a longer term opportunity, since pyrolysis oils are currently available only in small quantities. Howeve r, 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 ligni n can be converted to gasoline with reasonable H 2 

### consumption using mild hydrotreating follo wed by hydrocracking. The economics of this process are increasingly attractive with increased cr ude oil price or decreased pyrolysis oil price with increased production. At a price of $50/bbl fo r 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 pr ovided 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 CO 2 

### 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 CO 2 

### . 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 CO 2 

### emissions. Pyrolysis oil is less costly than natural gas by $1.2/MMBTU and therefore represents an economica lly viable substitute for natural gas in a refinery setting. The high oxygen content of pyrolysis oil im poses 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 a nd economic impact of hydrothermal treatment. Hydrotreating and hydrocracking, technologies wh ich were developed for oil refineries, were found to be particularly useful for processing vege table oil, greases, and pyrolysis oils. The level of hydrogen required for hydrotr eating can be controlled by pr ocessing conditions and catalyst, since the same materials which catalyze hydr odeoxygenation also catalyze decarboxylation. Optimization of this technology for biorefin eries will lead to better utilization of these biofeedstocks for conversion to fuels and chemicals. A key barrier to the use of biorenewables in o il refineries is that available quantities are small relative to the size of most refinery processing e quipment. Refiners are naturally resistant to risking catalysts and run time in large petrol eum 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 biorenewab les to be processed separately from petroleum fractions. These skid mounted units could also be used by agricultural and forest producers.



## 2

# II S TUDY 

# G OALS AND 

# B ASIS


## II.1 S TUDY 

## G OALS 

### The overall goal of this study was to evaluate the economics of using biomass-derived feedstocks (biofeedstocks) in petroleum refineries. The t echnical 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 evaluate d 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 modele d. Experiments were conducted as needed to determine the key processing parameters require d to understand cost and economic impact of biofeedstock use in a particular process. Cost est imates were developed and used in a broader evaluation of the potential process economics. Th is 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 u sed 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 gasolin e, 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 (Appendi x Tables 1 and 2) and additional background for the study basis. 

## II.2 S TUDY 

## B ASIS 

### 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 75050000

0 10000 20000 30000 40000



### Current Current Liquid Transport Fuels Liquid Transport Fuels Diesel Available Oil/Grease Available Oil/Grease



### Potential Potential Cellulosic Cellulosic Wastekbpd 

### 100

0 10 20 30 40 50 60 70 80 90



### Ethylene Ethylene Propylene Propylene Available Oil Available Oil

and Grease and GreaseM ton/yrM 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 Wastekbpd 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 Wastekbpd

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 Wastekbpd 

### 100

0 10 20 30 40 50 60 70 80 90



### Ethylene Ethylene Propylene Propylene Available Oil Available Oil

and Grease and GreaseM ton/yrM ton/yr 

### 100

0 10 20 30 40 50 60 70 80 90



### Ethylene Ethylene Propylene Propylene Available Oil Available Oil

and Grease and GreaseM ton/yrM ton/yr



### 0 10 20 30 40 50 60 70 80 90



### Ethylene Ethylene Propylene Propylene Available Oil Available Oil

and Grease and GreaseM ton/yrM 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 av ailable in the U.S. for conversion to fuels ( Table 1 ).Oils and grease are available today and are lik ely 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 th e potential supply is much larger. As shown in Figure 1 (a), significant levels of cellulosic wast e are available globally for production of pyrolysis oil if conversion to transportation fu els 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 W est Texas Intermediate (WTI) crude feedstock price of $40 per barrel. The impacts of price in creases 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 du ring 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, Fi gures 1 and 2). A transportation charge of $2 per barrel was assumed for biofeedstocks based on typical rail transportation costs for grease and oils.

• CO 2 

### credits. The use of pyrolysis oil as a fuel would generate a CO 2 

### credit which could be traded. Currently CO 2 

### is trading for $10/ton but is expected to reach $30/ton in 5
years.

• Subsidies . Potential subsidies were calculated b ased 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 project ed 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 expe nsive, 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 o il (VGO) to make gasoline or diesel.



## 5CH 4 H 2 H 2 Plant 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 i C 4

Reformer

CH 4 H 2 H 2 Plant 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 i C 4

Reformer



### Figure 4. Typical U.S. refinery.



### The ability to process biofeedstocks in existing re fineries is affected by their high oxygen and acid content. For example, oils and greases typical ly contain 11-12% oxygen while pyrolysis oils may contain 35-45% oxygen. The acidity of feeds pr ocessed 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 materi als 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 an alyses presented here assumed an average U.S. hydrogen plant supplying 25 million cubic feet per day of hydrogen using a natural gas feed.



## 6HCK SGP SGP

FCCU

VISB ISOM LPG Butanes

Gasoline

Jet

CSO Diesel

CH 4 Fuel Oil LCO NHT

MEROX C D U Crude

Oil

All H 2 Users MTBE

ALKY

F F F

H 2 Plant V D U Diesel Hydrotreating Reformer

HCK SGP SGP

FCCU

VISB ISOM LPG Butanes

Gasoline

Jet

CSO Diesel

CH 4 Fuel Oil LCO NHT

MEROX C D U Crude

Oil

All H 2 Users MTBE

ALKY

F F F

H 2 Plant V D U Diesel Hydrotreating Reformer



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



## 7

# III R EFINERY 

# O PPORTUNITIES FOR 

# V EGETABLE 

# O ILS AND 

# G REASE
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 di esel 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 overa ll 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 

### Feed Process Application 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 

### Feed Process Application Product

Veg oil/Grease



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



## III.1 B IODIESEL 

## (F ATTY 

## A CID 

## M ETHYL 

## E STER 

## ) P RODUCTION 

### 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 refini ng 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 re duce 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 sweeteni ng. Four alternatives are illustrated in Figure 7.

Figure 7. UOP Merox processing schemes.



### When a new hydrotreater is installed to lower th e amount of sulfur in a gasoline product, existing Merox equipment can become available for retrof it 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 esterificati on 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 requi res 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 a nother 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 worl d-wide biodiesel production rates increase, the reuse of crude glycerol as a refinery feedst ock 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 AcidExcess 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 AcidExcess 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 AcidExcess 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 C 4 

### olefins from fluid catalytic cracking units (FCCs) to methyl tertia ry 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 servi ce 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 Ethe rmax 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 exis ting 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 (<100 0 

### C). Because a typical FAME production unit has a re latively low processing capacity compared to most existing Merox or MTBE units, the unit tu rndown needs to be evaluated on a case-by-case basis. In principle, excess processi ng capacity can be judiciously used to provide longer reaction and/or separation times and to enhance hi gh one-pass conversions of triglycerides.
Preliminary Heterogeneous Catalyst Study 

### There are incentives for development of a hete rogeneous catalyst for FAME production whether it is employed in an existing or new unit:

• reduce operating costs (specifically consumption of chemicals)

• simplify product purification (eliminat e 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 prescr ibed 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, ex cess 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) s howed high one pass conversion of triglyceride to FAME at mild operating conditions (<100 0 

### 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 re quirements by 50% or more. Coupling the lower capital expenditures with renewable fuel subsid ies 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 (<100 0 

### C)

• an appropriate combination of feed pric e, 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 tes ting 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 li quid 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, th e conversion was much lower (less than 50%). Therefore further work to obtain a stable base ca talyst 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 G REEN 

## D IESEL 

## P RODUCTION 

### “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-Processi ng 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

H 2 2.9
Products
H 2 S + NH 3 0.04

C1 to C4 0.7 Naptha 2.4

CO 2 0.0

H 2 O 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 wh ile maintaining the total fresh feed rate. The crude-based feed rate is reduced to a pproximately 33100 bpd and supplemented with 1900 bpd brown grease. The product contains appr oximately 5% green diesel, as desired.

• Case II demonstrates the effect of increasi ng space velocity to accompany the additional biomass feed. Case II maintains the ratio betw een 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 b ase case. The treat-gas ratio is maintained at the base-case ratio, under the assumption that th e recycle gas compressor is not capacitylimited.

• Case III builds upon Case II by limiting the r ecycle 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 th e 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, consequent ly decreasing the catalyst temperature cycle. Capping the recycle gas compresso r 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 1T (°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 1.15R elative D e activationRate (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 32Nom 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 stai nless 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 industria lly-accepted limit before considering upgrades to the process unit metallurgy. Therefore, the react or 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 th e 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 possi ble 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-processi ng 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 hydrot reating 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 app ears 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 CO2 PSA

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 CO2 PSA

CH4 H2 Plant Feed
pump

Cooler



### Plant oils and yellow and brown greases may cont ain water, alkali metals (such as sodium, potassium), phosphorous, ash, and other contam inants. 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 catal yst 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, eith er 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 remova l mechanism from HDO to DeCO2. Decarboxylation and hydrodeoxygenation are illustrated in Figure 15

. C n COOH Pd / H C n + CO 2

C n COOH C n+1 + 2H 2 O +3H 2 Decarboxylation

Hydrodeoxygenation C n COOH Pd / H C n + CO 2

C n COOH C n+1 + 2H 2 O +3H 2 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

H 2 0.8 2.9

Products
H 2 S + NH 3 0.04 0.04

C1 to C4 0.6 0.7 Naptha 2.5 2.4

CO 2 15.5 0.0

H 2 O 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 CO 2 

### 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 CO 2 

### neutral, but actually produce additional CO 2 

### credits. Decarboxylation can not be cons idered in a unit producing ULSD for two reasons: (1) DeCO2 is favored over HDO at lower temperatures and pr essures than the conditions required for deep desulfurization of traditional feedstocks; (2) the production of CO 2 

### 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-al one unit, a pressure swing adsorption unit (PSA) system in the recycle gas stream can be provided to remove CO 2 

### . In this way, recycle gas purge, recycle gas and makeup gas compression are minimized.

The relevant biofeedstocks have low concentrati ons 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 hydr oprocessing 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 (h igher 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 estimat ed inside battery limits capital cost for a 2000 bpd standalone green diesel unit is appr oximately $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 estimat es could be generated from literature and modeling results alone. Our initial yield estimates assumed that hydrodeox ygenation would be the primary mechanism for oxygen removal, as reported by Canmet/Arbokem from their initial studies 14 

### run at 700-2200 psi and 350-450C.
However, experiments performed in this proj ect showed that both hydrodeoxygenation and decarboxylation occur, with the results summarized in Table 6 . Decarboxylation is favored at lower pressures and hydrodeoxygenation incr eased with increasing pressure. Decarboxylation results in odd number paraffin production and CO 2 

### 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 hydrot reating 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

% CO 2 + 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 H 2 

### gas is constantly flowing. For autoclave tests, losses are treated as gas which is calculated by diffe rence so gas yield is overestimated and liquid yield underestimated. Gas composition changed th roughout 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 presulf iding these materials. Therefore, there is considerable opportunity for improvement durin g development of a green diesel process, if pursued. 

## III.3 C OMPARISON OF 

## B IODIESEL AND 

## G REEN 

## D IESEL 

## P ROCESS AND 

## P ROPERTIES 

### 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 CO 2 

### . 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 G REEN 

## G ASOLINE AND 

## G REEN 

## O LEFINS 

## P RODUCTION 

### 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 cond itions 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
C 2 P+Methane 4.1 4.1

C 2 = 8.6 8.7

C 3 P 2.0 2.1

C 3 = 22.0 22.4

C 4 s 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
C 2 = 1.5 1.9

C 3 P 0.7 0.8

C 3 = 4.0 4.6

C 4 s 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 wo uld 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 ) wh ich 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 E CONOMIC 

## A SSESSMENT AND 

## L IFE 

## C YCLE 

## A NALYSIS 

### 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 averag e 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-viabl e feed for any of the products, because of its low cost. However, subsidies would have a signifi cant impact on the viability of all soy-based products, as can be seen by comparing the NPV w ith 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 ve getable 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 anal ysis 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 CO 2 

### ( Figure 20 ) than petroleum diesel. Of the biofuels, green diesel and green gasoline have the lowest environmental impacts and CO 2 

### production. The environmental impacts of biodiesel production are increased by the fact that it requires metha nol which is produced from natural gas. Methanol production is a very energy-intensive process so it brings with it a strong environmental burden.

0 5 0 0 0 5 0 0 0 5 0 0 0

Petroleum Diesel Biodiesel Green Diesel Green Gasoline 0 5 0 0 0 5 0 0 0 5 0 0 0

Petroleum Diesel Biodiesel Green Diesel Green GasolinePoints

0 5 0 0 0 5 0 0 0 5 0 0 0

Petroleum Diesel Biodiesel Green Diesel Green Gasoline 0 5 0 0 0 5 0 0 0 5 0 0 0

Petroleum Diesel Biodiesel Green Diesel Green GasolinePoints



### 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 Gasolinekg 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 Gasolinekg CO2/kg Fuel



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



## 24

# IV R EFINERY 

# O PPORTUNITIES FOR 

# P YROLYSIS 

# O IL
Pyrolysis oil, or bio-oil, is the material produ ced from pyrolysis of wood waste. Pyrolysis oil can be used as a substitute for fossil fuels to genera te heat, power, and/or chemicals. Transportation fuels such as methanol and Fischer-Tropsch fu els can be derived from pyrolysis oil through synthesis gas processes. Pyrolysi s 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 a lternatives 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 wate r-soluble components to produce hydrogen (Section IV.4 ) The study also investigated hydrothermal processi ng of pyrolysis oil and techniques for blending pyrolysis oil with other feedstocks (Section IV.5 ).


## IV.1 B ACKGROUND AND 

## P ROCESSING 

## O PTIONS 

### The properties of pyrolysis oil depend on the sp ecific feedstock and cond itions 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 flui dized 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 pyroly sis 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 Recover y 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.



## 26Pyrolysis 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 gua rd 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 w as 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 sta ndard 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 G ASOLINE 

## A ND 

## A ROMATICS 

## P RODUCTION FROM 

## P YROLYTIC 

## L IGNIN 

### In previous studies, whole pyrolysis oil was hydr otreated 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 leve l 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 consum ption. 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 ligni n 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, CO 2 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 copr ocessing 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 py rolytic 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/hydrocrack ing units have already been desi gned 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 C OPROCESSING 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 separ ate 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 pyrolysi s 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 consis tent 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)

C 2 - 2.0 3.3 3.6 3.8

C 3 P 1.2 2.1 2.4 .7

C 3 = 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 + CO 2 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 py rolytic lignin, the acid bio-oils appeared to increase the crackability of the VGO and shift VGO yiel ds toward increased light ends and lowerLCO 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 pyro lytic lignin and pyrolysis oil with VGO in an FCC might be possible if a separate feed system were used to inject the py rolysis oil or pyrolytic lignin. There are three concerns which would n eed 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 an d 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 a bout 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 th e 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 require d to minimize refiner risk before such an approach would be viable.



## 30

## IV.4 H2 PRODUCTION FROM 

## P YROLYSIS 

## O IL 

### 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 H 2 

### - 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 aci ds such as formic and acetic acids, with implications for processing unit metallurgy.

• Higher molecular weight carbohydrates such as sugars and anhydro-s ugars 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 tec hnology 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 (6 0500lb/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



## 31Aqueous 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 fo r the aqueous-phase of pyrolysis oil.



### methanation reactions occur in the adiabatic pre-re forming 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 produ ce the baseline quantity of hydrogen. The steam to feed ratio is varied to observe the effect on H 2 

### 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 f eed to the SMR reactor section can be completely displaced with the aqueous phase of the pyroly sis oil to produce at least the baseline H 2 

### generation rate of 25 MMscfd. The CO 2 

### emission per unit of hydrogen production is reduced by approximately 50 to 65% because components in th e aqueous phase of the pyrolysis oil do not contribute to CO 2 

### emissions while all carbon in the natu ral gas feed for SMR is eventually emitted as CO 2 

### .



## 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 CO 2 

### emission per unit of hydrogen is not sufficient to make up the difference through CO 2 

### 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 con cept 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 fro m 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 H 2 

### and reject a tail gas.

Table 22. Reforming equilibrium reactions for APR of pyrolysis oil components
Component Reforming Equilibrium Reaction Acetic Acid Reforming CH 3 (COOH) = 2 H 2 + 2 CO Acetol Reforming CH 3 (CO)CH 2 OH = 4 H 2 + 3 CO Ethylene Glycol Reforming H 2 C(OH)C(OH)H 2 = 3 H 2 + 2 CO

Formic Acid Reformi ng H(COOH) = H 2 O + CO Glyoxal Reforming H(CO)(CO)H = H 2 + 2 CO Hydroxyacetaldehyde Reforming H 2 C(OH)(CO)H = 2 H 2 + 2 CO Sugar Reforming C 6 H 12 O 6 = 6 H 2 + 6 CO

Water Gas Shift CO + H 2 O = H 2 + CO 2



### Process modeling suggests that part of the plant will be exposed to hot acetic vapor which poses metallurgical challenges. The co rrosivity of acetic acid solutions (and other organic acids like formic acid) is very dependent on process cond itions 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 ma y be viable alternatives, especially in oxygenfree environments, unless the material is transfe rred from tankage where oxygen content is a real concern. Development of a commercial APR pro cess will require corrosion testing at the various process conditions found throughout the process.



## 33Aqueous Phase of Pyrolysis Oil PSA H 2

Reactor Separator Tail Gas

Unconverted Pyrolysis Oil Water Recycle 

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



### The most conservative approach is to use titani um 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 approxim ately $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 pro cess 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 ar e 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 includ es acetic acid fractionation utilities. Acetic acid fractionation costs are approximately $0.06/lb H 2 

### generated. Table 23. Comparison of APR process flow schemes.
Process Concept COP ($/lb H 2 ) 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 routesto 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 H 2 

### 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 H YDROTHERMAL 

## P ROCESSING AND 

## B LENDING 

## S TUDIES 

### 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 st ep which could be done in the field to reduce the costs of transporting pyrolysis oil by rem oving oxygen and reducing acidity. Batch autoclave tests were performed to hydrothermally process th e 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 ke pt 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.Pyrolyticlignin,gResid gCatalystWater,gTemp CTime minPresPsi (max)Waterphase,gOrganicphaseGas,lGas,%CO2

1 75.7 0 75.1 300 15 138 0 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 75.0 350 120 248 0 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 20 280 0 77.0 46.5 3.3 47

8 75.1 75.3 300 15 210 0? 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 hydrot hermal 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 H 2 

### pressure breaks the molecules as the oxygen groups are removed, resulting in smaller molecules. Hydrothermal treat ment 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 op tion 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% Hyperm er 2296 and 1% Hypermer 256SF surfactants were used. Also encouraging was the fact that th e HT-82 sample of hydrotreated pyrolytic lignin was very soluble in both VGO and LCO.


## IV.6 E CONOMIC 

## A NALYSIS OF 

## P YROLYSIS 

## O IL 

## P ROCESSING 

## O PTIONS



### 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 SO 2 

### emissions. This would also allow refiners to reduce CO 2 

### emissions from boilers and fired heaters and obtain a CO 2 

### credit. This is an immediate appli cation for refiners and other industries using natural gas to minimize pollution. Minor cap ital investment would probably be needed to change burner tips and modify boilers to allow co-fi ring 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 substitu te 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 a queous 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 stea m 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 a pproximately 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 H 2 

### 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 H 2 

### production. The incentive is $3.5MM/year with the CO 2 

### credit and $0.3MM/year without. Table 29. Incentives for H 2 

### production
$/D Quantities Feed Water-soluble pyrolysis oil 52,500 5.250 bpd
Products Hydrogen 84,910 28 MMSCFD

CO 2 credit ($30/ton) 9,770

Utilities -31,500

Incentive 10,680 

### The capital cost of a pre-reformer to add to an existing H 2 

### 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


### CO 2 

### 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 O PPORTUNITIES FOR 

# F ISCHER 

# T ROPSCH 

# L IQUIDS

Another option for biorenewable use is gasifica tion 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 olefin s which are easily removed by hydrotreating. The study calculated the value of FT liquid to a typi cal refiner if biorefineries chose to send them to the refinery rather than upgr ade 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(av g) 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 inte gration 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 se nding the diesel and naphtha cut to DHT and the wax cut to HCK to produce diesel and naph tha. Capital is required for two towers and typical HCK conditions are not optimal for FT wax.

• Send FT liquids to their own tower, while se nding 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

CH 4 Fuel Oil LCO NHT

MEROX C D U Crude

Oil

All H 2 Users MTBE

ALKY

F F F

H 2 Plant V D U FT Liquid Diesel Hydrotreating Naphtha Reformer

HCK SGP SGP

FCCU

VISB ISOM LPG Butanes

Gasoline

Jet

CSO Diesel

CH 4 Fuel Oil LCO NHT

MEROX C D U Crude

Oil

All H 2 Users MTBE

ALKY

F F F

H 2 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

H 2 use -1.4 1200 -16.4 Operating cost and refinery margin -20

Total value $/ton 313



## 40

# VI A CRONYMS AND 

# A BBREVIATIONS



# . 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



## 41MMBtu 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 R EFERENCES

1 

### DOE/USDA. “Biomass as a Feedstock for a Bioe nergy and BioProducts Industry: The Technical Feasibility of a Billion-Ton Annual Supply”, April, 20 05, pp xii and 3. Joint study by DOE and USDA. Available at: http://www.eere.energy.gov/ biomass/pdfs/final_billi onton_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, “Bioma ss 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.barch art.com/comfund/tallow.asp

. 10 

### Spath, P. L., Dayton, D. C. (2003) Preliminar y 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 compou nds 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 Performa nce, 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 Incentiv e Fact Sheet. Available to members at: https://www.biodiesel.org/members/membersonly/f iles/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

. 20 

### Samolada MC and Vasolos IA. Catalytic Cracking of Bi omass Flash Pyrolysis Liquids” in Developments in Thermochemical Biomass Conversio n, 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 Techno logy” in Proceedings of the 9 th 

### European Bioenergy Conference 24-27 June 1996. Pergamon 1997, pp 1578-1583.

22 

### Bormann, Konrad; Tilgner, Hartmut. Influence of ra pe seed oil triglycerides on the catalytic cracking of petroleum distillate. Erdoel, Erdgas, Kohle (1994), 110(2), 75-7.



## 43

# O PPORTUNITIES FOR 

# B IORENEWABLES IN 

# O IL 

# R EFINERIES



# A PPENDIX 

# : B ACKGROUND AND 

# E XPERIMENT 

# S UMMARY

P REPARED 

# F OR 

# 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 inte rnational 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 th e 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 Biof eedstock 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