# PNNL-18284

Prepared for the U.S. Department of Energy

under Contract DE -AC05-76RL01830



# Production of Gasoline and Diesel

from Biomass via Fast Pyrolysis,

Hydrotreating and Hydrocracking:

A Design Case

SB Jones JE Holladay

C Valkenburg DJ Stevens

C W Walto n C Kinchin

DC Elliott S Czernik

February 2009

# PNNL-18284



# Production of Gasoline and Diesel

from Biomass via Fast Pyrolysis,

Hydrotreating and Hydrocracking:

A Design Case

SB Jone s 1 

# JE Holladay 1



# C Valkenburg 1 

# DJ Stevens 1



# C W Walton C Kinchin 2



# DC Elliott 1 

# S Czernik 2



# February 2009

Prepared for

the U.S. Department of Ener gy

under Contract DE-AC05-76RL01830

Pacific Northwest National Laboratory

Richland, Washington 99352

1 

# Pacific Northwest National Laboratory

2 

# National Renewable Energy Laboratory

# iii 

# Executive Summary

The President has established a goal to supply 35 billion gallons per year of renewable and alternative

fuels by 2017. This goal is addressed in part by the U.S. Department of Energy (DOE) Office of Biomass

Program‟s (OBP‟s) Thermochemical Platform multiyear program plan to “ convert biomass to fuels,

chemicals and power via thermal and chemical processes such as gasification, pyrolysis and other nonbiochemical processes ” ( DOE 2008 ).

In recent years, the Biomass Program completed technoeconomic evaluations of both biological and

thermochemical pathways for converting biomass to ethanol. These “design case” studies provided a

detailed basis for understanding the current state of various conversion technologies for producing fuel

ethanol. The studies also helped identify technical barriers for which research and development could

potentially lead to significant cost improvements. Consistent assumptions for items such as plant

lifetimes, rates of return, and other factors were used in all cases so the various processes could be

compared.

At present, the use of biomass resources to produce infrastructure-compatible fuels is appealing.

Hydrocarbon biofuels can potentially be used without significant changes to the current fuel distribution

and utilization infrastructure, including pipelines, pumping stations, and vehicles. Given the relatively

short time between now and 2017, the goal of 35 billion gallons per year of renewable fuels will be more

readily met if hydrocarbon biofuels are part of the fuel mix.

The purpose of this design case study is to evaluate a processing pathway for converting biomass into

infrastructure-compatible hydrocarbon biofuels. This design case investigates production of fast pyrolysis

oil from biomass and the upgrading of that bio-oil as a means for generating infrastructure-ready

renewable gasoline and diesel fuels. Other options for pyrolytic processes and upgrading steps exist, but

they were not evaluated in this study. Likewise, gasification pathways that could be used to produce

hydrocarbons are not addressed here. Th is study has been conducted using the same methodology and

underlying basis assumptions as the previous design cases for ethanol.

The overall concept and specific processing steps were selected because significant data on this

approach exists in the public literature. The analysis evaluates technology that has been demonstrated at

the laboratory scale or is in early stages of commercialization. The fast pyrolysis of biomass is already at

an early stage of commercialization, while upgrading bio-oil to transportation fuels has only been

demonstrated in the laboratory and at small engineeri ng development scale. Advanced methods of

pyrolysis, which are under development, are not evaluated in this study. These, may be the subject of

subsequent analysis by OBP.

The plant is designed to use 2000 dry metric tons/day of hybrid poplar wood chips to produce

76 million gallons/year of gasoline and diesel. The processing steps include:

1. Feed drying and size reduction

2. Fast pyrolysis to a highly oxygenated liquid product

3. Hydrotreating of the fast pyrolysis oil to a stable hydrocarbon oil with less than 2% oxygen

4. Hydrocracking of the heavy portion of the stable hydrocarbon oil

# iv

5. Distillation of the hydrotreated and hydrocracked oil into gasoline and diesel fuel blendstocks

6. Hydrogen production to support the hydrotreater reactors.

Note that the Idaho National Laboratory (INL) is working on feedstock logistics that will eliminate the

need for drying and size reduction at the plant. That is, the “as received” feedstock to the pyrolysis plant

will be “reactor ready”. This development will likely further decrease the cost of producing the fuel.

The capital cost for a standalone “n th 

# ” plant is $303 million (2007 basis). At a 10% return on

investment (ROI), the minimum fuels (gasoline + diesel) selling price is $2.04/gal ($1.34/gal ethanol

equivalent basis).

An important sensitivity is the possibility of co-locating the plant with an existing refinery. In this

case , the plant consists only of the first three steps: feed prep, fast pyrolysis, and upgrading. Stabilized,

upgraded pyrolysis oil is transferred to the refinery for separation and finishing into motor fuels. The offgas from the hydrotreaters is also transferred to the refinery, and in return the refinery provides lower-cost

hydrogen for the hydrotreaters. This reduces the capital investment to $188 million and the minimum

fuel selling price to $1.74/gal ($1.14/gal ethanol equivalent basis).

Production costs near $2/gal (in 2007 dollars) and petroleum industry infrastructure-ready products

make the production and upgrading of pyrolysis oil to hydrocarbon fuels an economically attractive

source of renewable fuels. The study also identifies technical areas where additional research can

potentially lead to further cost improvements.

# v 

# Acronyms and Abbreviations

AR As received

Btu British thermal units

CF B circulating fluid beds

DOE U.S. Department of Energy

gal gallon

HDS hydrodesulfurized

HTS high temperature shift

LHV lower heating value

MFSP minimum fuel product selling price

MM Million

mm millimeter(s)

mtpd metric tons per day

MWth megawatts thermal



# OBP Office of the Biomass Program

PSA pressure swing adsorption

ROI return on investment

tpd tons per day

TPEC total purchased equipment cost

# vii 

# Contents

Executive Summary .......................................................................................................................... iii

Acronyms and Abbreviations ............................................................................................................. v

1.0 Introduction ............................................................................................................................. 1.1

2.0 Analysis Approach .................................................................................................................. 2.1

3.0 Feedstock and Plant Size ......................................................................................................... 3.1

4.0 Process Overview .................................................................................................................... 4.1

4.1 Fast Pyrolysis Oil Upgrading .......................................................................................... 4.3

5.0 Process Design ......................................................................................................................... 5.4

5.1 Feed Handling and Fast Pyrolysis ................................................................................... 5.4

5.2 Hydrotreating to Stable Oil ............................................................................................. 5.6

5.3 Hydrocracking and Product Separation ........................................................................... 5.7

5.4 Hydrogen Production ...................................................................................................... 5.8

5.5 Cooling Water and Utilities ............................................................................................. 5.9

5.6 Process Yields and Consumptions ................................................................................ 5.10

6.0 Process Economics .................................................................................................................. 6.1

6.1 Capital Costs ................................................................................................................... 6.1

6.1.1 Pyrolysis Unit ....................................................................................................... 6.1

6.1.2 Pyrolysis Oil Hydrotreating and Hydrocracking .................................................. 6.1

6.1.3 Hydrogen Generation ........................................................................................... 6.2

6.1.4 Balance of Plant ................................................................................................... 6.2

6.2 Operating Costs ............................................................................................................... 6.3

6.3 Minimum Fuel Selling Price ........................................................................................... 6.3

7.0 Economic and Technical Sensitivities ..................................................................................... 7.1

7.1 Co -location with a Refinery ............................................................................................ 7.1

7.2 Financial and Market Sensitivities .................................................................................. 7.2

7.3 Research Sensitivities ...................................................................................................... 7.3

8.0 Conclusions and Recommendations ........................................................................................ 8.1

9.0 References ............................................................................................................................... 9.1

Appendix A Design Case

– Stand-Alone Heat and Material Balance ............................................ A.4

Appendix B Equipment Cost Details .............................................................................................. B.1

# viii 

# Figures

Figure 4.1. Block Diagram of Overall Design ................................................................................ 4.1

Figure 4.2. Typical Product Yields from Different Modes of Wood Pyrolysis .............................. 4.2

Figure 5.1. Block Diagram of Fast Pyrolysis .................................................................................. 5.4

Figure 5.2. Block Diagram of Bio-Oil Upgrading .......................................................................... 5.6

Figure 5.3. Block Diagram of Hydrocracking and Product Separation .......................................... 5.8

Figure 5.4. Block Diagram of Hydrogen Production ...................................................................... 5.9

Figure 7.1. Block Diagram of Pyrolysis and Upgrading Plant Co-located with a Refinery ........... 7.1

Figure 7.2. Financial/Market Sensitivities ...................................................................................... 7.3

Figure 7.3. Research Sensitivities ................................................................................................... 7.4



# Tables

Table 5.1. Feedstock and Processing Assumptions ........................................................................ 5.5

Table 5.2. Product Characterization ................................................................................................ 5.5

Table 5.3. Hydrotreating Conditions ............................................................................................... 5.7

Table 5.4. Hydrotreating Product Yields ........................................................................................ 5.7

Table 5.5. Hydrocracking Conditions ............................................................................................. 5.8

Table 5.6. Process Water Demands ................................................................................................ 5.9

Table 5.7. Overall Yields and Consumption ................................................................................. 5.10

Table 5.8. Carbon Balance ............................................................................................................ 5.10

Table 6.1. Pyrolysis Unit Literature Capital Cost .......................................................................... 6.1

Table 6.2. Total Project Investment Cost for the Design Case Stand-Alone Plant ......................... 6.2

Table 6.3. Total Project Investment Factors ................................................................................... 6.2

Table 6.4. Operating Cost Assumptions ......................................................................................... 6.3

Table 6.5. Economic Parameters ..................................................................................................... 6.4

Table 6.6. Project Economics for the Stand-Alone Design Case Plant........................................... 6.5

Table 7.1. Comparison of Integrated and Standalone Facilities ..................................................... 7.2

# 1.1 

# 1.0 Introduction

The President has established a goal to supply 35 billion gallons per year of renewable and alternative

fuels by 2017 (White House 2007). This goal is addressed in part by the U.S. Department of Energy

(DOE) Office of the Biomass Program‟s (OBP‟s) Thermochemical Platform multiyear program plan to

“ convert biomass to fuels, chemicals and power via thermal and chemical processes such as gasification,

pyrolysis, and other non-biochemical processes ” ( DOE 2008).

In recent years, the Biomass Program completed technoeconomic evaluations of both biological and

thermochemical pathways for converting bioma ss to ethanol. These “design case” studies provided a

detailed basis for understanding the current state of various conversion technologies for producing fuel

ethanol. The studies also helped identify technical barriers where research and development could

potentially lead to significant cost improvements. Consistent assumptions for items such as plant

lifetimes, rates of return, and other factors were used in all cases so the various processes could be

compared.

At present, the use of biomass resources to produce infrastructure-compatible fuels is appealing.

Hydrocarbon biofuels can potentially be used without significant changes to the current fuel distribution

and utilization infrastructure, including pipelines, pumping stations, and vehicles. Given the relatively

short time before 2017, the goal of 35 billion gallons of renewable fuels will be more readily met if

hydrocarbon biofuels are part of the fuel mix.

The purpose of this design case study is to evaluate a processing pathway for converting biomass into

infrastructure-compatible hydrocarbon biofuels.

This design case investigates fast pyrolysis oil production from biomass and the upgrading of that

bio-oil as a means for generating infrastructure-ready renewable gasoline and diesel fuels. The overall

concept and specific processing steps were selected because significant data on this approach exist in the

public literature. Other options for pyrolytic processes and upgrading steps exist, but they were not

evaluated in this study. One example of alternative processing options is hydrothermal pyrolysis

followed by other upgrading steps. Likewise, gasification pathways can be used to produce

hydrocarbons, but those are also not addressed here. These and other options may be addressed in futur e

studies.

The design case presented here represents a goal case targeting performance potentially available

between now and 2015. Th is analysis evaluates technology that has been demonstrated at the laboratory

scale or is in early stages of commercialization. The fast pyrolysis of biomass is already at an early stage

of commercialization, while the upgrading of the bio-oil to transportation fuels has only been

demonstrated in the laboratory and at small engineering development scale. As such, the analysis does

not reflect the current state of commercially-available technology but includes advancements that are

potentially achievable by 2015.

The study has been conducted using the same methodology and underlying basis assumptions as the

previous design cases for ethanol. It allows a basis for comparison with other research and development

projects targeting the DOE objectives and lastly, provides a benchmark for the status of the pyrolysis

program.

# 2.1 

# 2.0 Analysis Approach

The approach used is similar to that employed in previous conceptual process designs and associated

design reports (Aden et al. 2002, Spath et al. 2005, Phillips et al. 2007). Process flow diagrams are based

on literature information and research results; these data were then used to build a process model in

CHEMCAD © 

# , a commercial process flow sheeting software tool. The capital costs were taken from

literature sources or were obtained from Aspen ICARUS TM 

# software after being sized using the

CHEMCAD © 

# heat and material balances. The capital and operating costs were assembled in a Microsoft

Excel © 

# spreadsheet. A discounted cash flow method was used to determine the minimum product selling

price.

# 3.1 

# 3.0 Feedstock and Plant Size

The feedstock is hybrid poplar wood chips delivered at 50 wt% moisture. The feedstock cost is

assumed to be $50.7/bone dry metric ton, delivered. This report takes a conservative approach and

assumes the same moisture level as was used in previous studies (Aden et al. 2002, Spath et al. 2005,

Phillips et al. 2007) and that feedstock grinding and drying occur at the plant.

The plant capacity is 2000 metric tons/day (mtpd) of bone dry wood. The plant is assumed to be an

established (“n th 

# ”) plant design, rather than a first of its kind.

# 4.1 

# 4.0 Process Overview

A simplified block diagram of the overall design is given in Figure 4.1 . The processing steps include:

1. Feed drying and size reduction

2. Fast pyrolysis to a highly oxygenated liquid product

3. Hydrotreating of the fast pyrolysis oil to a stable hydrocarbon oil with less than 2% oxygen

4. Hydrocracking of the heavy portion of the stable hydrocarbon oil

5. Distillation of the hydrotreated and hydrocracked oil into gasoline and diesel fuel blendstocks

6. Steam reforming of the process off-gas and supplemental natural gas to produce hydrogen for the

hydrotreating and hydrocracking steps.

Figure 4.1 . Block Diagram of Overall Design

Feed Handling and Preparation: The biomass feedstock is dried from its as-received moisture

level to less than 10% to minimize water in the fast pyrolysis product liquid. It is then ground to 2-6 mm

particle size to yield sufficiently small particles, ensuring rapid reaction in the pyrolysis reactor.

Fast Pyrolysis: Pyrolysis is the thermal decomposition of carbonaceous material in the absence of

oxygen to produce char, gas, and a liquid product rich in oxygenated hydrocarbons. In general, pyrolysis

is performed using a range of temperatures and residence times to optimize the desired product.

Figure 4.2 illustrates the approximate yields from different modes of pyrolysis (Bridgewater 2007).

Fast pyrolysis is assumed in this work. The biomass is heated to approximately 500°C in less than 1

second, and then rapidly cooled to stop the reaction. The liquid product, known as bio-oil, is obtained in

yields up to 75% by weight on a dry feed basis. It can also be upgraded to lower the oxygen content and

transported using the same trillion-dollar infrastructure used by the oil industry. In addition to being

transported and stored at a lower cost than solid biomass, bio-oil and upgraded oil can be used in

applications ranging from value-added chemicals to transportation fuels.

# 4.2 Figure 4.2 . Typical Product Yields from Different Modes of Wood Pyrolysis

There are many types of pyrolysis reactor configurations, each with different advantages and

disadvantages. T he following paragraphs summarize the various reactor types:

Ablative pyrolysi s: Mechanical pressure is used to press carbonaceous material (usually wood)

against a heated reactor wall. Material in contact with the wall essentially “melts” and, as it is moved

away, the residual oil evaporates as pyrolysis vapors. There is no requirement for inert gas in this

configuration, so processing equipment (on demonstration scales) can be smaller and potentially less

expensive. The other advantage to this configuration is that feed materials do not require extensive

grinding to achieve particle sizes sufficient for high heating rates. Conversely, this configuration is

surface-area controlled so scaling is a linear function of the heat-transfer area and does not benefit

from the same economies of scale as the other systems. The system is also slightly more complex due

to the mechanical nature of the process. There is currently a 50 tpd demonstration plant using this

technology in Germany. (Bridgewater 2007)

Moving Bed or Auger Pyrolysis: This reactor utilizes a screw to mix hot sand and biomass. While it

provides good control of residence times and does not dilute the products with fluidizing gas, the sand

must be reheated separately, which leads to mechanical reliability issues. There are currently no

large-scale commercial plants. (Bridgewater 2007)

Entrained Flow Pyrolysis: This reactor configuration is popular for studies of thermochemical

conversion kinetics and investigations of the effects of pressure. Feed material is typically fed in to the

top of the reactor, co- cu rrent with a gas stream. The flow through this configuration is assumed to

approximate plug flow, with the residence time controlled by the length of the heated zone. Char

buildup can be a troublesome trade-off for these simple and inexpensive reactors. Liquid yields are

usually lower than fluid bed systems. (Bridgewater 2007)

Rotating Cone: This reactor combines biomass and hot sand at the top of the vessel. The solids are

mixed by a rotating cone inside the vessel. Hot pyrolysis oil vapors leave near the bottom of the

reactor, while hot sand and char exit the reactor from the bottom of the cone. No carrier gas is

needed, resulting in smaller downstream equipment. A 2 mtpd plant operating on palm oil empty fruit

bunches was commissioned in Malaysia in 2006. (BTG 2009)

Bubbling Fluidized Bed Pyrolysis: Here, biomass is introduced to hot sand fluidized by recirculated

product gas. This technology is well understood, simple to construct and operate (on large scales),

and very efficient in transferring heat to the biomass, resulting in high liquid yields. Small biomass

# 4.3

particle sizes are required for this method in order to obtain high heating rates. Dynamotive currently

operates two plants, one 100 mtpd and one 200 mtpd, in Canada. (Bridgewater 2007, Dynamotive

2007)

Circulating Fluidized Bed Pyrolysis : Similar to bubbling fluidized beds, these configurations circulate

hot sand between the pyrolysis reactor and a sand re-heater. Char or product gas and char are burned

to provide heat to the sand re-heater. Circulating fluid beds (CFBs) are widely used at very high

throughputs in the petroleum industry and are potentially suitable for large biomass throughputs as

well. Ensyn operates a 100 tpd (200 tpd green wood) plant in Canada. (Bridgewater 2007, Ensyn

2009)

Both fluidized bed configurations allow for greater control of vapor residence time, an important

consideration for optimization of the high liquid yield associated with fast pyrolysis. A circulating

fluidized bed was chosen, as this design is the most likely to be scalable to the 2000 mtpd feed rate

assumed in this report.



# 4.1 Fast Pyrolysis Oil Upgrading

Untreated bio-oil is a dark brown, free-flowing liquid with about 25% water that cannot be easily

separated. It is a complex mixture of oxygenated compounds that is unstable in long-term storage

(Oasmaa and Kuoppala 2003, Diebold 2000 ) and is not miscible with any conventional hydrocarbonbased fuel (Bridgewater 2007). Bio-oil can be stabilized and converted to a conventional hydrocarbon

fuel by removing the oxygen through hydrotreating. Hydrotreating to remove nitrogen and sulfur from

hydrocarbons is a common and well established refinery process. Oxygen removal on the scale needed to

upgrade pyrolysis oil is relatively new and is in the research stages.

The upgrading step involves contacting the bio-oil with hydrogen under pressure and at moderate

temperatures (<400°C or 750°F) over fixed bed reactors. Single stage-hydrotreating has proved to be

difficult, producing a heavy, tar-like product. Dual-stage processing, where mild hydrotreating is

followed by more severe hydrotreating has been found to overcome the reactivity of the bio-oil (Elliot

2007). Overall, the pyrolysis oil is almost completely deoxygenated by a combination of

hydrodeoxygenation and decarboxylation:

Catalyst / 3H 2



# C

n 

# COOH C n+1 

# + 2H 2 

# O Hydrodeoxygenation

Catalyst / H 2



# C

n 

# COOH C

n 

# + CO 2 

# Decarboxylation

Less than 2% oxygen remains in the treated, stable oil. Water and off-gas are produced as

byproducts. The water phase contains some dissolved organics, while the off-gas contains light

hydrocarbons, excess hydrogen, and carbon dioxide.

Once stabilized oil is produced it can be further processed into conventional fuels or sent to a

refinery. Both options are considered in this report.

# 5.4 

# 5.0 Process Design

The design case assumes a standalone unit that combines feed pretreatment, fast pyrolysis,

hydrotreating and hydrocracking of pyrolysis oil and separation to gasoline and diesel fuel blendstocks,

and hydrogen generation.



# 5.1 Feed Handling and Fast Pyrolysis

Figure 5.1 shows the block diagram for the feed pretreatment, fast pyrolysis, and bio-oil quenching

area. Hybrid poplar is dried from its „as received (AR)‟ moisture level of 50 wt% to 7 wt% water using

hot flue gases from the char combustor in a direct-contact dryer. It is then ground to approximately 2-6

mm

.

Figure 5.1 . Block Diagram of Fast Pyrolysis

Table 5.1 summarizes the as received feedstock characteristics and feed requirements.

The dried, finely ground biomass is fed to a circulating fluidized pyrolysis reactor operating at 520°C.

Sand is used as the fluidizing medium and the residence time is less than a second. The biomass is

converted into a mixture of gases, bio-oil, and char. A cyclone separates the sand and char from the gases

and liquids. The hot bio-oil vapor is rapidly quenched with cooled bio-oil and then separated from the

remaining vapors (Solantausta 2003, Freel and Graham 1995). High temperature heat recovery is not

included due to the likelihood of severe fouling in the recuperators (Johnson et al. 2006). Most of the

gases are recycled back to the pyrolysis reactor to assist fluidization. The char and a portion of the gas

are burned to heat the circulating sand. The cooled pyrolysis oil contains about 20-25% associated water.

# 5.5

Table 5.2 shows the product characterization. Note that the bio-oil yield is near the high end of

the references. This choice is based on the assumption that improvements in bio-oil yield can be achieved

by the 2015 time-frame.

Table 5.1 . Feedstock and Processing Assumptions

Assumed Value

Hybrid Poplar Wood Chips

AR Moisture, wt% 50%

wt % Carbon (dry) 50.60 (a)



# wt % Hydrogen (dry) 6.08 (a)



# wt % Oxygen (dry) 40.75 (a)



# wt % Nitrogen (dry) 0.61 (a)



# wt % Sulfur (dry) 0.02 (a)



# wt % Chlorine ( dry) 0.01 (a)



# wt % Ash (dry) 1.93 (a)



# HHV

, B tu

/lb 8405 (a)



# Temperature, °F (°C) 968 (520) (b)



# Pressure atmospheric (b)



# Feed Moisture to pyrolysis unit, wt% 7 (b)



# (a) Hybrid Poplar. Phyllis Database for Biomass and Waste

(b) Solantausta 2003



# Table 5.2 . Product Characterization

Model Results Reference Data (a) 

# Reference Data (b)



# Yields, lb/100 lb dry wood

Oil 65 59.9 66

Water 1 0 10.8 12

Char & Ash 1 3 16.2 8

Gas 1 2 13.1 11

Loss 3

Oil Composition

Water in oil, wt% 2 1 15

- 30

Carbon, wt% dry 58 55

- 58

Hydrogen, wt% dry 6 5.5

- 7.0

Oxygen, wt% dry 36 35

- 40

(a) Ringer et al. 2006

(b) Mohan et al. 2006

# 5.6

The detailed model flowsheet diagram and associated heat and material balance for the fast pyrolysis

section can be found in Appendix A.



# 5.2 Hydrotreating to Stable Oil

The simplified flow diagram for the stand-alone upgrading portion of the plant is shown in Figure 5.2

.

Figure 5.2 . Block Diagram of Bio-Oil Upgrading

The filtered bio

- oil product from the pyrolysis unit is pumped to high pressure, then combined with

compressed hydrogen and preheated with reactor effluent. Two catalytic reaction stages are used. The

first stage ca talytic reactor serves to stabilize the bio

- oil by mild hydrotreatment over cobalt molybdenum

( CoMo ) hydrotreating catalyst (Elliott 2007) . The product oil is further processed in the second

- stage

hydrotreater

. The second stage hydrotreater operates at h igher temperature and lower space velocity than

the first stage. CoMo catalyst is also used in this reactor. Reactor effluent and a fired heater provide feed

preheat for the second stage . The second

- stage product is separated into product oil, wastewate r

, and off

gas streams

.

The off

- gas from the hydrotreaters is sent to a Pressure Swing Adsorption ( PSA ) system for recovery

of the hydr ogen gas. The recovered hydrogen is recycled back to the reactors. The low pressure PSA tail

gas stream

, which is rich in by product light hydrocarbons, is sent to the steam reformer for hydrogen

production. A small portion of this off

- gas is used in the fired heater that preheats feed for the second

stage hydrotreater

.

Anaerobic digestion can be used to treat the less th an 2% organics in the wastewater

.

The processing conditions are listed in Table 5.3 . The product oil is a mixture of hydrocarbons with a

low level (~2%) of oxygen, as shown in Table 5.4 . Data from Beckman et al. (1990) and Elliot (2007)

were used to estimate the product slate.

# 5.7 Table 5.3

. Hydrotreating Conditions

1 st 

# Stage (a) 

# 2 nd 

# Stage (a)



# Temperature, °F (°C) 4 65 (2 4 0) 700 (3 70 )

Pressure, psig 2500 2015

LHSV, v/h/v 1 0.14

(a) Beckman et al 1990



# Table 5.4

. Hydrotreating Product Yields

Model Reference

Data ( a ) 

# Reference

Data ( b )



# Yields, lb /100 lb wet pyrolysis oil

Stable Oil (Stream 304) 4 4 38

Water (Stream 230) 48 50

Gas (Streams 270 and 302) 1 3 12 by difference

Chemical H 2 

# Consumption, lb/100 lb dry oil 4. 96 5.01 3.45

Stable Oil Composition (Stream 304)

Water, ppm 0 50 0

Aromatics, wt% 10.0 Not reported

Carbon, wt% dry 88. 1 86.8 86.8

Hydrogen, wt% dry 1 0.5 13.2 10.8

Oxygen, wt% dry 1.5 0.02 2.5

Specific Gravity 0.8 7 0.83 0.93

B tu /lb, gross 17, 6 00 19,765 17,302

B tu /lb, net 16, 6 00 18,525 16,276

( a ) Tables 15 & 16 in Elliot 2007

( b ) Beckman et al. 1990 , Chapter 5



# 5.3 Hydrocracking and Product Separation

Figure 5.3 shows the simplified flow diagram for the hydrocracking and product separations areas.

The hydrotreated oil is stabilized by removing the butane and lighter components in a lights removal

column. The stable oil stream is then separated into light and heavy fractions. The heavy fraction (which

boils above 350°C (662°F) is sent to the hydrocracker to completely convert the oil to gasoline and diesel

blend components.

# 5.8 Figure 5.3 . Block Diagram of Hydrocracking and Product Separation

The product is a mixture of liquids spanning the gasoline and diesel range and some byproduct gas.

The gasoline and diesel range products are separated by distillation. These products are suitable for

blending into finished fuel. The hydrocracking conditions and product yields are shown in Table 5.5

.

Table 5.5

. Hydrocracking Conditions

Mode l Reference Data

Nominal Temperature, °F (°C) 800 (427) 775 (413) (a)



# Pressure, psig 1280 1051 (mild)

– 2400 (resid) (a)



# Gasoline, lb/100 lb feed 9 5 Depends on feed/conditions (a,b)



# Diesel, lb/100 lb feed 5 Depends on feed/conditions (a,b)



# (a) Parkash 2 003

(b) Gulf 2006



# 5.4 Hydrogen Production

Figure 5.4 shows the simplified flow scheme for hydrogen generation by steam reforming of natural

gas (SRI International 2007, Meyers 2004) and the off-gas streams from the hydrotreating and

hydrocracking processes. The off-gas by itself is insufficient to produce the amount of hydrogen required

by the hydrotreaters and hydrocrackers; supplemental natural gas is required. Most of the off-gas is used

to fire the reformer; however, a portion of the off-gas is compressed and mixed with makeup natural gas

which is then hydrodesulfurized (HDS). Hydrogen for the HDS unit is supplied by the off-gas stream.

The gas exiting the HDS is then mixed with superheated steam and sent to the steam reformer to produce

syngas. The syngas hydrogen content is increased by high temperature shift (HTS). After condensing out

the water, the hydrogen is purified by pressure swing adsorption (PSA). Off-gas from the PSA is

recycled to the reformer burners.

# 5.9 Figure 5.4 . Block Diagram of Hydrogen Production

Saturated and superheated steam are generated by recuperating heat from the reformer exhaust and

during syngas cooling. The steam is used to provide steam to the reformer and heat for process heaters

and distillation column reboilers. A steam export credit is taken for the small amount of excess steam not

used by the process. The heat and material balance for the steam reformer can be found in Appendix A.



# 5.5 Cooling Water and Utilities

Cooling water usage is minimized through the use of air fin coolers where applicable. The process

water demands are shown in Table 5.6 . Most of the cooling tower water is used to indirectly cool the fast

p yr olysis bio-oil that is recirculated in the quench system. Optimizing the quench process could reduce

this value further. For example, chillers could supply some of the cooling, however this would increase

the power demand. Electricity is bought from the grid; no power is generated onsite. Waste water is

treated in an onsite facility before discharge.

Table 5.6

. Process Water Demands

Source of Water Demand Model

Cooling tower makeup, gpm 373

Steam reformer boiler feed water makeup, gpm 113

Total, gpm 486

Overall demand, gal water/gal product 3.0

# 5.10



# 5.6 Process Yields and Consumptions

The yield of product fuel from wood is shown in Table 5.7 . The unadjusted yield, shown in the first

ro w of Table 5.7, does not take into account the natural gas used in steam reforming for hydrogen

production. It is of interest to know what the yield of fuel products would be if only renewable feedstocks

were used. In the base case, hydrogen for the upgrading reactors is generated in a conventional steam

reformer using offgas from the upgrading reactors plus natural gas. This hydrogen could also be derived

from renewable resources, for example by 1) gasifying biomass, then shifting the syngas to hydrogen or

by 2) steam reforming a portion of the fast pyrolysis oil. It is estimated that if the natural gas used to

generate hydrogen was replaced by biomass or bio-oil, approximately one third of the biomass feedstock

would be needed for hydrogen generation. The yield of fuel products on a 100% renewable basis is

shown in the second row of Table 5.7. The third row of Table 5.7 shows the 100% renewable yield on an

ethanol equivalent basis (lower heating value (LHV) adjusted). Table 5.7 also summarizes water, power

and NG consumption.

Table 5.7

. Overall Yields and Consumption

Model

Fuel Product, gal/st dry wood

– unadjusted for natural gas usage in the reforme r 100

Fuel Product, gal/st dry wood

– adjusted to 100% renewable feedstock basis 6 5

LHV ethanol equivalent Fuel Product, gal/st dry wood , 100% renewable basis

~ 100

Gallons makeup water/gallons fuel product 3.0

Power, kWh/gallon fuel product 2 .5

Natura l Gas, scf/gallon fuel product 42



# Table 5.8 shows the carbon balance. More than half of the carbon in the combined biomass plus

natural gas feed is converted into fuel products.

Table 5.8

. Carbon Balance

Feeds Biomass 88 %

Natural Gas 1 2 %

sum 100%

Fuel Products Gasoline Pool 23 %

Diesel Pool 32 %

Fuels sum 5 5 %

Waste Products Pyrolysis Unit Exhaust 23 %

Upgrading Wastewater 0%

Upgrading Heaters Exhaust 2 %

Reformer Exhaust 20%

Waste sum 4 5 %



# The heat and material balance for the Stand-Alone Design Case is located in Appendix A.

# 6.1 

# 6.0 Process Economics

6.1 Capital Costs

The cost basis assumes that this is the n th 

# plant. Capital costs for the pyrolysis unit, the hydrocracker

unit, and the hydrogen plant are from literature sources. The heat and material balance was used to size

standard equipment using CHEMCAD © 

# sizing routines or by using standard literature sources (Perry et

al. 1984, Couper et al. 2005). ASPEN Icarus TM 

# was used to obtain bare equipment costs for sized

equipment. The 2007 annual Chemical Engineering Index was used to escalate capital costs to 2007.

Equipment sizing and cost sources are summarized below. Additional details can be found in Appendix

B.



# 6.1.1 Pyrolysis Unit

Existing pyrolysis units process 200 tpd of biomass or less, and capital cost data is limited. This

report used data from Ringer et al. 2006, scaled to 2000 tpd. Costs associated with feed handling,

pyrolysis reactor, quench, and product recovery were included. Heat recovery and power generation were

excluded. The 35% equipment contingency used in Ringer et al. 2006 was retained. As seen in Table

6.1 , these costs are within the average of the data shown.

Table 6.1

. Pyrolysis Unit Literature Capital Cost

Study

Date Original Scale,

mtpd 2007 TIC MM $

at 2000 tpd Power,

kW

- h /tonne Source Notes

2003 96 68.7 40 (a) 

# Cost for 5 96

- tpd units

2003 490 55.5 156 (b) 

# Estimated from 40

- 60 MWth data

2007 200 48.5 (c) 

# Dynamotive Plant

2006 550 44.5 192 (d) 

# TIC includes 35% equipment contingency

1990 1001 33.7 (e) 

# 20 fluidized beds @ 100 tpd level

2003 500 39.2 156 (b) 

# Estimated from 15

- 140 MWth data

2003 96 35.5 (f) 

# Cost scaled with 6/10 rule

1990 1001 27.1 167 (e) 

# 1 circulating fluidized

- bed pyrolysis unit

(a) Bridgewater et al 2002

(b) Solantausta 2003

(c) Dynamotive 2007

(d) Ringer et al 2006

(e) Beckman et al 1990

(f) Johnson et al 2006



# 6.1.2 Pyrolysis Oil Hydrotreating and Hydrocracking

Most of the equipment for this section were sized using CHEMCAD © 

# and costed using ASPEN

Icarus TM 

# or taken from vendor budgetary quotes. The cost of a small hydrocracker system processing

hydrotreated bio-oil has been estimated by UOP (Marker et al. 2005). The large-diameter, high-pressure

hydrotreater vessels are likely to be manufactured outside of the United States. A 15% equipment

contingency was added to account for uncertainty in the fabrication and hydrocracker unit costs.

# 6.2



# 6.1.3 Hydrogen Generation

Hydrogen generation by steam reforming of natural gas is a well-known, mature process; therefore no

equipment contingency was added. The SRI Yearbook (SRI International 2007) and Foster Wheeler

(Gulf 2006, Meyer 2004) have published capital and operating data for plants of various capacities.



# 6.1.4 Balance of Plant

Storage capacity and cooling tower capital were estimated with ASPEN Icarus. Wastewater

treatment costs are from Beckman et al. 1990. A 15% process contingency was added.

Table 6.2 and Table 6.3 list the stand-alone design case installed equipment capital costs and the

factors used to determine the total project investment, respectively. The factors in Table 6.2 are from

Peters et al (2003).

Table 6.2

. Total Project Investment Cost for the Design Case Stand-Alone Plant

2007 MM USD % Contribution

Fast Pyrolysis 92 30 %

Hydrotreating 81 2 7 %

Hydrocracking and Separations 2 9 10 %

Hydrogen Generation 86 28 %

Utilities, etc. 14 5 %

303* 100

* Summation higher due to rounding



# Table 6.3

. Total Project Investment Factors

Total Purchased Equipment Cost (TPEC) 100%

Purchased Equipment Installation 39% of TPEC

Instrumentation and Controls 26% of TPEC

Piping 31% of TPEC

Electrical Systems 10% of TPEC

Buildings (including services) 29% of TPEC

Yard Improvements 12% of TPEC

Total Installed Cost (TIC) 2 47%

Indirect Costs

Engineering 32% of TPEC

Construction 34% of TPEC

Legal and Contractors Fees 23% of TPEC

Project Contingency 37% of TPEC

Total Indirect 126%

Total Project Investment 373%

# 6.3



# 6.2 Operating Costs

Table 6.4 lists the assumptions used to calculate the operating costs. Catalyst costs were escalated

with the average Producer Price Index for 2007.

Table 6.4

. Operating Cost Assumptions

Value

R aw Materials

Hybrid Poplar Chips, $/dry short ton 50.70

Ash Disposal, $/short ton 18 (a)



# Hydrogen (refinery transfer), $/lb 0.56 (b)



# Hydrotreating Catalyst , $/lb (200 7 ) 1 5. 5 ( c )



# Hydrocracking Catalyst, $/lb (200 7 ) 15. 5 ( c )



# Hydrogen Plant Catalysts, $/1 000 scf H 2 

# (200 7 ) 3.6 ( c )



# Utilities

Natural Gas, $/1000scf (1000 B tu /scf) 7. 68 ( d )



# Electricity, ¢ /kwh 6.3 6 ( d )



# Labor

Operating labor, $/hr burdened & 10% shift overlap 37.66 ( c )



# Maintenance and o verhead 95 % of l abor & s upervision (e)



# Materials

Maintenance 2 % of t otal project

investment ( e )



# Local taxes & insurance 2% of total project investment e )



# (a) Phillips et al. 2007

(b) Holmgren et al. 2007

(c) SRI International 200 7

(d) Energy Information Agency 2007

(e) Phillips et al. 200 7



# 6.3 Minimum Fuel Selling Price

The minimum fuel product selling price (MFSP) for the gasoline and diesel blendstock was

determined using a discounted cash flow rate of return analysis. The methodology is identical to that used

in Phillips et al. (2007). The MFSP is the selling price of the fuel that makes the net present value of the

process equal to zero with a 10% discounted cash flow rate of return over a 20 year plant life. The stream

factor (90%) is lower than that used in Phillips et al. (2007) to account for the shorter catalyst life (1 year

versus 5) assumed in this study. While two products are produced, (motor gasoline blendstock and diesel

# 6.4

blendstock), they are combined and referred to as a “fuel product” for simplicity. All MFSP calculations

are performed and reported on a combined product basis. Table 6.5 gives the economic parameters used

to calculate the MFSP. A sensitivity analysis was conducted to determine the effect of different financial

and operating assumptions on the MFSP.

Table 6.5 . Economic Parameters

Value

Stream Factor 90%

MACRS Depreciation, yrs 7 (a)



# Internal Rate of Return, % 10% (a)



# Plant life, yrs 20 (a)



# Construction Period

1 st 

# 6 months ‟ expenditure

Next 12 months ‟ expenditur e

Last 12 months ‟ expenditure 2.5 years (a)



# 8% (a)



# 60% (a)



# 32% (a)



# Start

- up time

Revenues

Variable Costs

Fixed Costs 6 months (a)



# 50% (a)



# 75% (a)



# 100% (a)



# Working Capital 5% of Total Capital Investment (a)



# Land 6% of Total Purchased Equipment Cost

(t aken as 1 st 

# year construction expense) (a)



# (a) Phillips et al. 2007



# The variable operating costs were determined from the heat and material balances and checked

against published data for the hydrocracker and steam reformer. These and the fixed operating costs are

shown in Table 6.6

.

# 6.5 Table 6.6

. Project Economics for the Stand-Alone Design Case Plant

$ /gal product Contribution

Feedstock 0.48 2 3 %

Natural Gas 0.32 16%

Catalysts & Ch emicals 0.15 7 %

Waste Disposal 0.01 0 %

Utilities (Cooling Water

, Electricity, Steam) 0.17 8%

Fixed Costs (Labor, Operating Supplies, etc.) 0.22 11 %

Capital Depreciation 0.20 1 0 %

Average Income Tax 0.13 7%

Average ROI 0.36 18 %

MFSP, $/gal 2.0 4

MFSP Ethanol Equivalent Basis, $/gal 1.3 4



# The 2007 average refiner prices for gasoline and diesel were $2.18 and $2.20 respectively (Energy

Information Agency 2007). Thus, at $2.04/gal fuel product price, there is incentive to pursue motor fuels

fr om biomass.

# 7.1



# 7.0 Economic and Technical Sensitivities

The design case describes a single operating point for a stand-alone processing unit. This section

investigates the production cost sensitivities to technical, financial, and market issues. These includ e

plant size, ROI , feedstock costs, reactor conditions and product yields. Because hydrocarbon fuels are the

final product, an important sensitivity is the possibility of co-locating the pyrolysis and upgrading facility

with an existing refinery.



# 7.1 Co -location with a Refinery

An alternate configuration to the standalone plant described in Section 5 is a facility that is co-located

wi th an existing refinery. Figure 7.1 shows the design case integrated with a refinery. The pyrolysis unit

is the same as in the stand-alone case. Co-location eliminates the need for a PSA unit in the hydrotreating

section if the upgrading unit off-gas can be sent to refinery hydrogen generation. In return, the upgrading

unit receives refinery hydrogen at a lower cost. All final processing of the stable oil to fuels occurs in the

refinery.

Figure 7.1 . Block Diagram of Pyrolysis and Upgrading Plant Co-located with a Refinery

VTT and partners (Solantuasta 2006) and UOP (Marker et al 2005, Holmgren et al 2007) are

separately pursuing producing bioproducts compatible with existing refinery infrastructure. This will

lessen the initial investment; however, direct use of stable oil in a refinery requires compl ete

deoxygenation and a low acid number to prevent corrosion. As an alternative, a finishing deoxygenation

unit and/or a separate hydrotreating/hydrocracking operation using dedicated stainless steel equipment

may be needed

. T he finishing hydrotreating u nit may result in saturating aromatic compounds, in which

case the light naphtha portion may need to be reformed in the refinery into aromatics before blending into

gasoline

.

# 7.2

Table 7.1 compares the stand-alone and integrated economics. The capital cost calculations assume

that the refinery equipment is fully depreciated and only operating costs apply to the refinery side. This

greatly reduces the capital investment. A hydrogen transfer price of $0.56/lb (Holmgren et al 2007)

assumes that all operating costs (including feedstock and fuel) to the hydrogen generation unit within the

refinery are accounted for in that price. A credit is taken for the fuel gas sent from the upgrading unit to

the refinery. The transfer price for the fuel gas to the refinery is determined by adjusting the current price

of natural gas for the Btu value of the fuel gas. The product costs are significantly lower than those for

the stand-alone plant due to reduced capital investment and the net cost difference between hydrogen and

fuel gas.

Table 7.1 . Comparison of Integrated and Standalone Facilities

Capital

Investment,

MM $ Minimum Selling

Price $/gal fuel

product Minimum Selling

Price $/gal ethanol

equiv alent

Stand

- A lone Plant

– Design

Case 3 03 2.04 1.3 4

Integrated Plant

– Design Case 188 1. 74 1.14



# 7.2 Financial and Market Sensitivities

The sensitivities to financial assumptions are shown in Figure 7.2 . Refinery integration with a

$0.56/lb hydrogen transfer price reduces the base case MFSP by about $0.30/gal. No refining capital

costs were assumed. However, operating costs based on the amount of stable oil processed were

included. It is possible that some revamping of refinery equipment will be needed to process the stable

oil. The cost of modifying a hydrocracker was estimated by assuming that the incurred costs would not

be greater than a small hydrocracker dedicated to stable oil. At a $0.56/lb hydrogen transfer price, this

option reduces the base MFSP by about $0.13/gal. Finally, the hydrogen transfer price in an integrated

facility was varied by $0.75 /lb and $1.25/lb hydrogen. The MFSP of the integrated facility exceeds that

of the base case at hydrogen transfer prices above $1.00/lb, which is the approximate transfer price for

hydrogen if new hydrogen generation facilities must be built.

Sensitivities to project contingency, capital investment, scale and ROI were determined for the standalone case. Reducing the project contingency to three percent of the total capital project, (similar to that

used in Philli ps et al. 2007), reduces the selling price by about $0.22/gal. The accuracy of the capital

estimate is in the range of -10% to +40%, and gives an MFSP range of $1.79/gal to $2.22/gal or a

difference of -0.22/gal to +$0.20/gal.

The base plant assumes a 90% stream factor. This may be optimistic with single reactors for fast

pyrolysis and upgrading. The equipment spare sensitivity determines the product cost increase if both the

pyrolysis unit and upgrading units had 3 reactors at 50% capacity each. This increases the selling price

by about $0.21/gal.

The cost of a single small pyrolysis with an upgrading unit at 500 tpd increases the selling price to

almost $2.68 /g al. Four 500 tpd units feeding a single upgrading unit is more cost effective, assuming the

feedstock cost is the same for both cases. The cost of transporting biomass feedstock is a limiting factor

# 7.3

in the size of a biomass processing plant. It is possible that small, distributed pyrolysis plants can take

advantage of lower feedstock costs and produce an easily transported product that can then be upgraded to

fuel in a large, centralized facility. Further analysis is needed to look at the tradeoff between pyrolysis

plant size and feedstock cost, and the transportation costs to a centralized upgrading plant.

Figure 7.2 . Financial/Market Sensitivities

The base case assumes a 10% ROI . A 30% return (shown on a separate scale to prevent making the

other sensitivities illegible) nearly doubles the baseline estimated selling price. The 0% return represents

the plant gate costs.



# 7.3 Research Sensitivities

The research sensitivities are shown in Figure 7.3 . Feedstock cost, reactor space velocity, reactor

pressure, and catalyst life all impact the minimum selling price for the standalone case. None of these

costs alone make the process uneconomic. However, the potential combined effects, such as catalyst life

and space velocity, are significant enough to warrant further research. Catalyst costs are important

because there is little data on life and the appropriate catalyst composition. As a base case, this study

assumes standard hydrotreating and hydrocracking catalysts.

The base case assumes that the two stage hydrotreaters described in Section 5.2 sufficiently

deoxygenate the stable oil, and only the heavy tail end requires hydrocracking. If this is not sufficient,

# 7.4

and the entire stable oil stream is sent to the hydrocracker, the size of the hydrocracker increases

significantly. Hydrocracking the entire stable oil stream adds $0.10/gal to price of the fuel.

Figure 7.3 . Research Sensitivities

# 8.1



# 8.0 Conclusions and Recommendations

Th is analysis suggests that production of hydrocarbon motor fuels from biomass via the pyrolysis

route is potentially economically attractive. The cost becomes even more favorable if the facility can be

closely associated with an existing petroleum refinery to leverage that infrastructure. Both the standalone plant that combines fast pyrolysis with hydrotreating, upgrading and hydrocracking to fuels and the

refinery-integrated plant fit with the existing infrastructure in terms of products and processing methods.

Developing partnerships with industry suppliers and refiners is important for production development and

product acceptance.

The pyrolysis pathway for producing hydrocarbon fuels is economically competitive with either the

biological or thermochemical pathways for production of ethanol fuels. On an energy-equivalent basis, as

shown in Table 7.1 , the costs of either hydrocarbon or alcohol fuels are essentially the same when

consistent analysis methodologies and underlying financial assumptions are used.

As indicated previously, other pyrolytic processes exist, and the economics of those processes may be

different from the one analyzed here. This analysis suggests that cost-effective processing routes to

hydrocarbon fuels exist, but does not attempt to compare this approach with others. If other pyrolytic

processing routes have lower costs, the production of hydrocarbon fuels will be even more attractive.

The analysis describes a system with performance projected to be available in 2012. Further research

and development is needed to understand the limitations of the process and where improvements can be

made. The sensitivity analysis (Figure 7.3) suggests that the biggest impacts of technical improvements

are in the area of catalysis as it relates to bio-oil upgrading. Catalyst lifetimes and performance are

essential to the process and need additional development and testing. Pyrolytic processing of the biomass

is a relatively small percentage of the overall cost, and changes to the pyrolysis process will have little

direct impact on the fuel cost. However, research on pyrolytic processing that leads to a better quality

bio-oil requiring less upgrading could also potentially reduce product costs.

The analysis (Figure 7.2) also shows the fuel costs are highly sensitive to plant size and ROI. The

close coupling of the upgrading steps with a petroleum refinery helps leverage the economies of scale in

the fuel production steps. The pyrolysis step to produce bio-oil is less sensitive to scale and can

potentially be decoupled from the fuel upgrading step. This analysis uses figures for ROI that are

consistent with previous biofuels analyses produced by DOE and allow comparison with other technology

options. Industries attempting to build actual facilities may use different assumptions, and the costs for

the resulting fuels would change accordingly.

To achieve the potential for producing hydrocarbon fuels from biomass, additional research and

development are needed. A summary of needs includes the following:

Understand the trade-off between the size of the pyrolysis oil plant and feedstock and transportation

costs to a centralized upgrading facility.

Incorporate the reductions in feed preprocessing (grinding and drying) by using a reactor ready feed.

Understand pyrolysis reactor and upgrading reactor scale-up limitations if any.

Improve the quality and consistency of bio-oil.

# 8.2

Determine detailed characterization of upgraded oil and products, for example:

– boiling point curves and densities

– compound types (% aromatics, naphthenes, olefins, iso and normal) within the gasoline, diesel

and heavy fraction ranges

– acid number.

Conduct catalyst life studies for the pyrolysis oil upgrading catalysts.

Develop better understanding of reactor limitations and opportunities by developing a predictive

pyrolysis-oil production model and upgrading the reactor model to replace the stoichiometric models

used in this study.

Determine stable oil and product fuel specifications for oil refinery acceptance.

In addition, there are new initiatives for an advanced pyrolysis process for biomass, the details of

which fall outside this base-case study. Incorporation of those new concepts into economic assessments

of future cases for biomass pyrolysis should be the next stage of analysis. Advanced pyrolysis processes

may include:

new reactor system engineering involving improved heat transfer mechanisms

pyrolysis including in -situ catalysis

pyrolysis including in -situ chemical processing

post pyrolysis processing to improve product bio-oil properties

in -process separation systems to produce bio-oil fractions with useful properties

post-process separations and subsequent treatment to produce improved products

higher efficiency integrated systems with improved heat utilization.

# 9.1 

# 9.0 References

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# 9.2

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Parkash, S. 2003. “Hydrocracking Processes.” Chapter 3 in Refining Processes Handbook . Gulf

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# ed. McGrawHill, New York.

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# A.4 

# Appendix A

Design Case

– Stand-Alone Heat and Material BalanceA.1



# Figure A.1 . Flow Diagram for Pyrolysis UnitA.2 

# Table A.1 . Pyrolysis Unit Heat and Material Balance

Stream No. 100 105 106 111 113 116 120 126 130 131 133

Temp C 20 176.587 66 499.421 499.421 592 15.5555 1211.84 499.421 43.7381 43.746

Pres psig 3 3 3 1 1 1 0.3041 3 1 1 10

Enth MMBtu/h

- 1743

- 1324.9

- 575.9

- 6192.3

- 5449.1

- 5322.2

- 1.4752

- 157.77

- 743.17

- 96574

- 96573

Vapor mass fraction 0 1 0 1 1 0 1 1 1 0 4.17E

- 05

Total lb/h 367436 560708 197546 3135535 2779270 2755767 350000 390817 356265 2.7E+07 2.7E+07

Flowrates in lb/h

Oxygen 0 24433.8 0 3873.85 3873.85 0 81520.1 24435 0 0 0

Nitrogen 0 269595 0 106 73.8 0 0 268480 269595 10673.8 80.5111 80.496

Water 183718 180076 13828.2 32280.7 0 0 0 10186 32280.7 4480534 4480506

Hybrid Poplar 183718 0 183718 0.0041 0.0041 0 0 0 0 0 0

Hydrogen 0 2.0727 0 1403.29 963.918 0 0 2.0729 439.369 1.3593 1.359

Carbon 0 0 0 15065.9 15065.9 0 0 0 0 0 0

Carbon Monoxide 0 513.829 0 175437 0 0 0 513.888 175437 1272.26 1272.03

Carbon Dioxide 0 86088.1 0 9738.7 0 0 0 86085.5 9738.7 2331.92 2331.05

Methane 0 0.0002 0 876.195 0 0 0 0.0002 876.195 27.2094 27.2046

Ethylene 0 0.0 004 0 3488.06 0 0 0 0.0004 3488.06 475.291 475.223

Ethane 0 0 0 0 0 0 0 0 0 0 0

Propylene 0 0.0006 0 3629.54 0 0 0 0.0006 3629.54 2174.05 2173.91

Propane 0 0 0 0 0 0 0 0 0 0 0

I

- Butane 0 0 0 0 0 0 0 0 0 0 0

Ammonia 0 0.0003 0 285.508 0 0 0 0.0003 285. 508 468.22 468.193

Pyro

- lignin 0 0 0 64485 0 0 0 0 64485 1.2E+07 1.2E+07

Cellobiose 0 0 0 22689.1 0 0 0 0 22689.1 4500798 4501477

Levoglucosan 0 0 0 5970.83 0 0 0 0 5970.83 1087896 1087571

Furfural 0 0 0 11941.7 0 0 0 0 11941.7 2078346 2078007

Hydroxy Acetone 0 0 0 5970.83 0 0 0 0 5970.83 1098279 1098360

Acetic Acid 0 0 0 8359.16 0 0 0 0 8359.16 1192790 1192921

Formaldehyde 0 0 0 0 0 0 0 0 0 0 0

2

- 5

- Xylenol 0 0 0 0 0 0 0 0 0 0 0

N

- Heptane 0 0 0 0 0 0 0 0 0 0 0

1

- ts

- 35

- 3C1CycC6 0 0 0 0 0 0 0 0 0 0 0

3

- 3

- 5

- TriMth

- C7 0 0 0 0 0 0 0 0 0 0 0

N

- PropylCyc

- C6 0 0 0 0 0 0 0 0 0 0 0

1

- 2

- 3

- Mesitylene 0 0 0 0 0 0 0 0 0 0 0

N

- ButylCycHexane 0 0 0 0 0 0 0 0 0 0 0

1

- 2

- DiC1

- 3C2

- Bz 0 0 0 0 0 0 0 0 0 0 0

Cis

- Decalin 0 0 0 0 0 0 0 0 0 0 0

Dimethyl

- C11 0 0 0 0 0 0 0 0 0 0 0

1

- 2

- 4

- triethylbe 0 0 0 0 0 0 0 0 0 0 0

1

- 1

- Bicyclohexyl 0 0 0 0 0 0 0 0 0 0 0

Diphenyl 0 0 0 0 0 0 0 0 0 0 0

diamantane 0 0 0 0 0 0 0 0 0 0 0

Phenanthrene 0 0 0 0 0 0 0 0 0 0 0

N

- C15

- CycloC5 0 0 0 0 0 0 0 0 0 0 0

Hexatriacontane 0 0 0 0 0 0 0 0 0 0 0

Ca 0 0 0 3496.15 3496.15 0 0 0 0 0 0

Sulphur 0 0 0 36.0271 36.0271 0 0 0 0 0 0

Calcium Chloride 0 0 0 67.057 67.057 0 0 0 0 0 0

Sand 0 0 0 2755767 2755767 2755767 0 0 0 0 0A.3 

# Table A.1 . (contd)

Stream No. 137 138 139 140 144 145 146 147 148 150 161

Temp C 43.746 37.7778 43.746 43.7381 43.7381 43.7381 91.8145 91.8145 91.8145 43.7199 105.294

Pres psig 10 5 10 1 1 1 8 8 8 1 3

Enth MMBtu/h

- 96072

- 96222

- 501.34

- 391.99

- 69.9

- 332.32

- 327.7

- 33.491

- 294.2

- 571.24

- 37.313

Vapor mass fra ction 4.17E

- 05 4.70E

- 05 4.17E

- 05 1 0 1 1 1 1 5.38E

- 05 0.50092

Total lb/h 3E+07 2.7E+07 139274 216984 12363 204621 204621 20912.8 183708 151637 9599.24

Flowrates in lb/h

Oxygen 0 0 0 0 0 0 0 0 0 0 0

Nitrogen 80.078 80.0782 0.4179 10673.4 0 10 673.4 10673.4 1090.85 9582.52 0.4179 0

Water 4E+06 4457246 23259.8 8992.01 8902.09 89.9201 89.9201 9.1901 80.73 32161.8 6000

Hybrid Poplar 0 0 0 0 0 0 0 0 0 0 0

Hydrogen 1.352 1.352 0.0071 439.362 0 439.362 439.362 44.904 394.458 0.0071 0

Carbon 0 0 0 0 0 0 0 0 0 0 0.0002

Carbon Monoxide 1265.4 1265.43 6.6035 175430 0 175430 175430 17929.4 157500 6.6035 0

Carbon Dioxide 2319 2318.95 12.1013 9725.74 0 9725.74 9725.74 993.998 8731.74 12.1012 0

Methane 27.063 27.0633 0.1412 876.049 0 876.049 876.049 89. 5347 786.514 0.1412 0

Ethylene 472.76 472.756 2.467 3485.52 0 3485.52 3485.52 356.23 3129.29 2.467 0

Ethane 0 0 0 0 0 0 0 0 0 0 0

Propylene 2162.6 2162.62 11.2855 3618.11 0 3618.11 3618.11 369.781 3248.33 11.2855 0

Propane 0 0 0 0 0 0 0 0 0 0 0

I

- Buta ne 0 0 0 0 0 0 0 0 0 0 0

Ammonia 465.76 465.763 2.4305 283.051 0 283.051 283.051 28.9286 254.123 2.4305 0

Pyro

- lignin 1E+07 1.2E+07 64281.4 0.044 0.044 0 0 0 0 64281.4 0

Cellobiose 4E+06 4478109 23368.6 0.0129 0.0129 0 0 0 0 23368.7 0

Levoglucosan 1E+0 6 1081925 5645.94 0.0011 0.0011 0 0 0 0 5645.94 0

Furfural 2E+06 2067219 10787.6 814.207 814.207 0 0 0 0 11601.8 0

HydroxyAcetone 1E+06 1092658 5701.95 349.864 349.864 0 0 0 0 6051.81 0

Acetic Acid 1E+06 1186728 6192.84 2296.82 2296.82 0 0 0 0 8489.66 0

Formaldehyde 0 0 0 0 0 0 0 0 0 0 0

2

- 5

- Xylenol 0 0 0 0 0 0 0 0 0 0 0

N

- Heptane 0 0 0 0 0 0 0 0 0 0 0

1

- ts

- 35

- 3C1CycC6 0 0 0 0 0 0 0 0 0 0 0

3

- 3

- 5

- TriMth

- C7 0 0 0 0 0 0 0 0 0 0 0

N

- PropylCyc

- C6 0 0 0 0 0 0 0 0 0 0 0

1

- 2

- 3

- Mesitylene 0 0 0 0 0 0 0 0 0 0 0

N

- ButylCycHexane 0 0 0 0 0 0 0 0 0 0 0

1

- 2

- DiC1

- 3C2

- Bz 0 0 0 0 0 0 0 0 0 0 0

Cis

- Decalin 0 0 0 0 0 0 0 0 0 0 0

Dimethyl

- C11 0 0 0 0 0 0 0 0 0 0 0

1

- 2

- 4

- triethylbe 0 0 0 0 0 0 0 0 0 0 0

1

- 1

- Bicyclohexyl 0 0 0 0 0 0 0 0 0 0 0

Diphenyl 0 0 0 0 0 0 0 0 0 0 0

diamantane 0 0 0 0 0 0 0 0 0 0 0

Phenanthrene 0 0 0 0 0 0 0 0 0 0 0

N

- C15

- CycloC5 0 0 0 0 0 0 0 0 0 0 0

Hexatriacontane 0 0 0 0 0 0 0 0 0 0 0

Ca 0 0 0 0 0 0 0 0 0 0 3496.16

Sulphur 0 0 0 0 0 0 0 0 0 0 36.0271

Calcium Chloride 0 0 0 0 0 0 0 0 0 0 67.057

Sand 0 0 0 0 0 0 0 0 0 0 0A.4 

# Table A.1 . (contd)

Stream No. 1810 1811

Temp C 32.2222 40.5555

Pres psig 50 45

Enth MMBtu/h

- 68430

- 68279

Vapor mass fraction 0 0

Total lb/h 10044315 10044310

Flowrates in lb/h

Oxygen 0 0

Nitroge n 0 0

Water 10044315 10044310

Hybrid Poplar 0 0

Hydrogen 0 0

Carbon 0 0

Carbon Monoxide 0 0

Carbon Dioxide 0 0

Methane 0 0

Ethylene 0 0

Ethane 0 0

Propylene 0 0

Propane 0 0

I

- Butane 0 0

Ammonia 0 0

Pyro

- lignin 0 0

Cellobiose 0 0

Levoglucosa n 0 0

Furfural 0 0

HydroxyAcetone 0 0

Acetic Acid 0 0

Formaldehyde 0 0

2

- 5

- Xylenol 0 0

N

- Heptane 0 0

1

- ts

- 35

- 3C1CycC6 0 0

3

- 3

- 5

- TriMth

- C7 0 0

N

- PropylCyc

- C6 0 0

1

- 2

- 3

- Mesitylene 0 0

N

- ButylCycHexane 0 0

1

- 2

- DiC1

- 3C2

- Bz 0 0

Cis

- Decalin 0 0

Dim ethyl

- C11 0 0

1

- 2

- 4

- triethylbe 0 0

1

- 1

- Bicyclohexyl 0 0

Diphenyl 0 0

diamantane 0 0

Phenanthrene 0 0

N

- C15

- CycloC5 0 0

Hexatriacontane 0 0

Ca 0 0

Sulphur 0 0

Calcium Chloride 0 0

Sand 0 0A.5



# Figure A.2 . Flow Diagram for Pyrolysis Oil Stabilization

# A.6 Table A.2 . Pyrolysis Oil Hydrotreating Heat and Material Balance

Stream No. 200 202 203 205 207 208 212 213 215 220 221

Temp F 110.6798 174.8342 430 465.5045 470 700 484.8971 481.4931 319.3935 150 110

Pres psig 1 2500 2495 2495 2490 2475 2475 2470 710.3041 705.3041 700.3041

Enth MMBtu/h

- 579.38

- 572.13

- 533.75

- 533.75

- 532.53

- 493.17

- 493.17

- 494.39

- 532.77

- 560.91

- 566.01

Vapor mass fraction 0 0.062243 0.15953 0.24881 0.25969 0.52647 0.6898 0.67505 0.22969 0.12881 0.12332

Total lb/h 151607.6 161371 161371 161371 161371 161371 161371 161374.8 161374.8 161374.8 161374.8

Flowrates in lb/h

Water 32161.65 32161.64 32161.64 44375.22 4 4375.22 44375.22 71936 71936 71936 71936 71936

Hydrogen 0 9763.431 9763.431 8474.411 8474.411 8474.411 3839.811 3843.556 3843.557 3843.557 3843.557

Carbon Dioxide 0 0 0 1000 1000 1000 9000.002 9000.002 9000.003 9000.003 9000.004

Methane 0 0 0 50 50 50 3 350 3350 3350 3350 3350

Ethane 0 0 0 6.9 6.9 6.9 2006.899 2006.9 2006.9 2006.9 2006.9

Propane 0 0 0 5.1 5.1 5.1 1705.1 1705.101 1705.101 1705.101 1705.101

I

- Butane 0 0 0 4.4 4.4 4.4 1504.4 1504.4 1504.4 1504.4 1504.4

Pyro

- lignin 64305.6 64305.58 64305. 58 54819.78 54819.78 54819.79 0 0 0 0 0

Cellobiose 23313.63 23313.62 23313.62 13988.17 13988.17 13988.17 0.0037 0.0051 0.0051 0.0051 0.0051

Levoglucosan 5674.065 5674.063 5674.063 2724.021 2724.021 2724.021 0.001 0.0014 0.0014 0.0014 0.0014

Furfural 116 33.4 11633.4 11633.4 0 0 0 0 0 0 0 0

HydroxyAcetone 6044.922 6044.922 6044.922 4835.938 4835.938 4835.938 0 0 0 0 0

Acetic Acid 8474.371 8474.367 8474.367 5084.619 5084.619 5084.619 0 0 0 0 0

2

- 5

- Xylenol 0 0 0 26002.45 26002.45 26002.45 8000.004 8000.00 2 8000.001 8000.001 8000.001

N

- Heptane 0 0 0 0 0 0 3275.653 3275.653 3275.654 3275.654 3275.655

1

- ts

- 35

- 3C1CycC6 0 0 0 0 0 0 4892.768 4892.769 4892.768 4892.768 4892.769

3

- 3

- 5

- TriMth

- C7 0 0 0 0 0 0 1994.707 1994.707 1994.707 1994.707 1994.707

N

- PropylC yc

- C6 0 0 0 0 0 0 4892.767 4892.769 4892.769 4892.769 4892.77

1

- 2

- 3

- Mesitylene 0 0 0 0 0 0 642.0663 642.0664 642.0664 642.0664 642.0665

N

- ButylCycHexane 0 0 0 0 0 0 234.2028 234.2029 234.2029 234.2029 234.2029

1

- 2

- DiC1

- 3C2

- Bz 0 0 0 0 0 0 1593.217 1593.2 17 1593.217 1593.217 1593.217

Cis

- Decalin 0 0 0 0 0 0 3186.433 3186.433 3186.433 3186.433 3186.433

Dimethyl

- C11 0 0 0 0 0 0 8635.319 8635.321 8635.321 8635.321 8635.322

1

- 2

- 4

- triethylbe 0 0 0 0 0 0 3186.432 3186.434 3186.434 3186.434 3186.433

1

- 1

- Bicyc lohexyl 0 0 0 0 0 0 238.9825 238.9825 238.9825 238.9825 238.9825

Diphenyl 0 0 0 0 0 0 4383.289 4383.289 4383.289 4383.289 4383.289

diamantane 0 0 0 0 0 0 8635.319 8635.321 8635.321 8635.321 8635.322

Phenanthrene 0 0 0 0 0 0 6087.9 6087.901 6087.901 6087 .901 6087.902

N

- C15

- CycloC5 0 0 0 0 0 0 291.442 291.4421 291.442 291.442 291.442

Hexatriacontane 0 0 0 0 0 0 0 0 0 0 0

Chrysene 0 0 0 0 0 0 6120.283 6120.284 6120.283 6120.283 6120.283

Cyclopentane 0 0 0 0 0 0 0 0 0 0 0

N

- Pentadecane 0 0 0 0 0 0 0 0 0 0 0

N

- Octadecane 0 0 0 0 0 0 0 0 0 0 0

Tetralin 0 0 0 0 0 0 0 0 0 0 0

Ethylbenzene 0 0 0 0 0 0 0 0 0 0 0

Benzene 0 0 0 0 0 0 0 0 0 0 0

N

- Butane 0 0 0 0 0 0 0 0 0 0 0

Toluene 0 0 0 0 0 0 0 0 0 0 0

1

- 3

- 5

- Mesitylene 0 0 0 0 0 0 0 0 0 0 0

P

- Xylene 0 0 0 0 0 0 1738.065 1738.064 1738.064 1738.064 1738.065

N

- Propylbenzene 0 0 0 0 0 0 0.0131 0.0102 0.0102 0.0102 0.0102

# A.7 Table A.2 . (contd)

Stream No. 223 225 226 227 230 240 242 245 250 253 258

Temp F 100 99.9329 99.9329 200 100 100 44.3597 100 77 84 .6947 303.1303

Pres psig 700.3041 35.3041 35.3041 35.3041 700.3041 700.3041 35 300 300 300 2500

Enth MMBtu/h

- 33.374

- 0.07606

- 33.298

- 30.008

- 487.88

- 46.01

- 45.02

- 0.20059

- 0.90996

- 1.1106 6.2841

Vapor mass fraction 0 1 0 0.000449 0 1 0.99656 1 1 1 1

Total lb/h 69520.25 25.2322 69495.02 69495.01 72156.03 19698.48 16435.05 3263.436 6500 9763.431 9763.431

Flowrates in lb/h

Water 242.2865 0.2879 241.9986 241.9986 71621.36 72.3108 72.3111 0 0 0 0

Hydrogen 0.4663 0.455 0.0113 0.0113 3.7543 3 839.336 575.9001 3263.436 6500 9763.431 9763.431

Carbon Dioxide 31.1286 17.3243 13.8043 13.8043 250.624 8718.251 8718.251 0 0 0 0

Methane 0.6464 0.6206 0.0258 0.0258 5.2046 3344.148 3344.149 0 0 0 0

Ethane 0.5571 0.5241 0.0331 0.0331 0.0423 2006.3 2006. 3 0 0 0 0

Propane 730.0195 3.4583 726.5612 726.5612 17.2753 957.8062 957.8062 0 0 0 0

I

- Butane 889.0693 2.1131 886.9562 886.9562 6.5343 608.7963 608.7963 0 0 0 0

Pyro

- lignin 0 0 0 0 0 0 0 0 0 0 0

Cellobiose 0 0 0 0 0.0051 0 0 0 0 0 0

Levoglucosan 0 0 0 0 0.0014 0 0 0 0 0 0

Furfural 0 0 0 0 0 0 0 0 0 0 0

HydroxyAcetone 0 0 0 0 0 0 0 0 0 0 0

Acetic Acid 0 0 0 0 0 0 0 0 0 0 0

2

- 5

- Xylenol 7748.411 0.0064 7748.406 7748.406 249.3144 2.2761 2.2761 0 0 0 0

N

- Heptane 3193.759 0.2507 3193.508 3193.508 0.712 8 81.1817 81.1817 0 0 0 0

1

- ts

- 35

- 3C1CycC6 4857.789 0.0991 4857.69 4857.689 0.1593 34.8195 34.8196 0 0 0 0

3

- 3

- 5

- TriMth

- C7 1987.447 0.0195 1987.427 1987.427 0.015 7.2457 7.2457 0 0 0 0

N

- PropylCyc

- C6 4878.879 0.0392 4878.839 4878.839 0.1611 13.7284 13.7 284 0 0 0 0

1

- 2

- 3

- Mesitylene 641.6517 0.001 641.6508 641.6508 0.0785 0.3363 0.3363 0 0 0 0

N

- ButylCycHexane 233.9581 0.0007 233.9575 233.9575 0.0024 0.2423 0.2423 0 0 0 0

1

- 2

- DiC1

- 3C2

- Bz 1592.735 0.0012 1592.733 1592.733 0.0573 0.4248 0.4248 0 0 0 0

Ci s

- Decalin 3184.489 0.0052 3184.483 3184.483 0.0499 1.8948 1.8948 0 0 0 0

Dimethyl

- C11 8633.4 0.0045 8633.395 8633.395 0.0018 1.9188 1.9189 0 0 0 0

1

- 2

- 4

- triethylbe 3186.106 0.0008 3186.105 3186.105 0.0099 0.3178 0.3178 0 0 0 0

1

- 1

- Bicyclohexyl 238.9565 0.0001 238.9565 238.9565 0.0004 0.0256 0.0256 0 0 0 0

Diphenyl 4383.134 0.0003 4383.133 4383.133 0.0454 0.1092 0.1092 0 0 0 0

diamantane 8633.234 0.0054 8633.229 8633.229 0.0045 2.0826 2.0826 0 0 0 0

Phenanthrene 6087.885 0 6087.884 6087.884 0.0107 0.00 52 0.0052 0 0 0 0

N

- C15

- CycloC5 291.442 0 291.442 291.442 0 0.0001 0.0001 0 0 0 0

Hexatriacontane 0 0 0 0 0 0 0 0 0 0 0

Chrysene 6120.283 0 6120.283 6120.283 0.0003 0 0 0 0 0 0

Cyclopentane 0 0 0 0 0 0 0 0 0 0 0

N

- Pentadecane 0 0 0 0 0 0 0 0 0 0 0

N

- Octadecane 0 0 0 0 0 0 0 0 0 0 0

Tetralin 0 0 0 0 0 0 0 0 0 0 0

Ethylbenzene 0 0 0 0 0 0 0 0 0 0 0

Benzene 0 0 0 0 0 0 0 0 0 0 0

N

- Butane 0 0 0 0 0 0 0 0 0 0 0

Toluene 0 0 0 0 0 0 0 0 0 0 0

1

- 3

- 5

- Mesitylene 0 0 0 0 0 0 0 0 0 0 0

P

- Xylene 1732.51 0.0 15 1732.494 1732.494 0.6275 4.927 4.927 0 0 0 0

N

- Propylbenzene 0.0102 0 0.0102 0.0102 0 0 0 0 0 0 0

# A.8 Table A.2 . (contd)

Stream No. 270 736 738 1820 1821

Temp F 44.4289 700.0002 496.55 90 105

Pres psig 35 645 645 60 55

Enth MMBtu/h

- 45.096

- 22.4 68

- 25.758

- 2320.7

- 2315.6

Vapor mass fraction 0.99656 1 1.00E

- 06 0 0

Total lb/h 16460.28 4037.127 4037.127 341022.1 341022.1

Flowrates in lb/h

Water 72.5986 4037.127 4037.127 341022.1 341022.1

Hydrogen 576.3552 0 0 0 0

Carbon Dioxide 8735.575 0 0 0 0

Methane 3344.769 0 0 0 0

Ethane 2006.824 0 0 0 0

Propane 961.2645 0 0 0 0

I

- Butane 610.9094 0 0 0 0

Pyro

- lignin 0 0 0 0 0

Cellobiose 0 0 0 0 0

Levoglucosan 0 0 0 0 0

Furfural 0 0 0 0 0

HydroxyAcetone 0 0 0 0 0

Acetic Acid 0 0 0 0 0

2

- 5

- Xy lenol 2.2824 0 0 0 0

N

- Heptane 81.4324 0 0 0 0

1

- ts

- 35

- 3C1CycC6 34.9186 0 0 0 0

3

- 3

- 5

- TriMth

- C7 7.2652 0 0 0 0

N

- PropylCyc

- C6 13.7676 0 0 0 0

1

- 2

- 3

- Mesitylene 0.3373 0 0 0 0

N

- ButylCycHexane 0.243 0 0 0 0

1

- 2

- DiC1

- 3C2

- Bz 0.426 0 0 0 0

Cis

- Decalin 1 .9 0 0 0 0

Dimethyl

- C11 1.9233 0 0 0 0

1

- 2

- 4

- triethylbe 0.3186 0 0 0 0

1

- 1

- Bicyclohexyl 0.0257 0 0 0 0

Diphenyl 0.1094 0 0 0 0

diamantane 2.088 0 0 0 0

Phenanthrene 0.0052 0 0 0 0

N

- C15

- CycloC5 0.0001 0 0 0 0

Hexatriacontane 0 0 0 0 0

Chrysene 0 0 0 0 0

Cyclopentane 0 0 0 0 0

N

- Pentadecane 0 0 0 0 0

N

- Octadecane 0 0 0 0 0

Tetralin 0 0 0 0 0

Ethylbenzene 0 0 0 0 0

Benzene 0 0 0 0 0

N

- Butane 0 0 0 0 0

Toluene 0 0 0 0 0

1

- 3

- 5

- Mesitylene 0 0 0 0 0

P

- Xylene 4.942 0 0 0 0

N

- Propylbenzene 0 0 0 0 0

# A.9 Figure A.3 . Flow Diagram for Hydrocracking and Product Separation

# A.10 Table A.3 . Hydrocracking and Product Separation Heat and Material Balance

S tream No. 227 302 304 305 306 307 308 309 310 500 508

Temp F 200 203.3611 457.477 308.4405 427.7641 470 734.4869 398.6693 398.8038 744.7188 366

Pres psig 35.3041 35.3041 35.3041 5 5 2 1 1 15 1300 1295

Enth MMBtu/h

- 30.008

- 4.1857

- 15.689

- 12.8

- 5.614 5

- 1.3824 5.6133

- 10.415

- 10.413 5.6658 5.651

Vapor mass fraction 0.000449 1 0 0 0 0.52048 0 0 0 0 0.20116

Total lb/h 69495.01 3466.51 66028.52 19360.17 46668.37 46668.38 8800.003 37868.36 37868.37 8800.005 11010

Flowrates in lb/h

Water 241. 9986 241.9984 0 0 0 0 0 0 0 0 0

Hybrid Poplar Ch 0 0 0 0 0 0 0 0 0 0 0

Hydrogen 0.0113 0.0113 0 0 0 0 0 0 0 0 2129.363

Carbon Dioxide 13.8043 13.8043 0 0 0 0 0 0 0 0 0

Methane 0.0258 0.0258 0 0 0 0 0 0 0 0 0

Ethane 0.0331 0.0331 0 0 0 0 0 0 0 0 0

Pro pane 726.5612 726.5609 0 0 0 0 0 0 0 0 0

I

- Butane 886.9562 886.956 0 0 0 0 0 0 0 0 0

2

- 5

- Xylenol 7748.406 0.0001 7748.407 3162.434 4585.977 4585.977 0.0083 4585.968 4585.968 0.0083 0.0083

N

- Heptane 3193.508 1596.752 1596.765 1596.766 0 0 0 0 0 0 0

1

- ts

- 35

- 3C1CycC6 4857.689 0.0486 4857.643 4541.519 316.1233 316.1233 0.0016 316.1216 316.1216 0.0016 0.0016

3

- 3

- 5

- TriMth

- C7 1987.427 0.0002 1987.428 1856.982 130.4458 130.4458 0.0003 130.4456 130.4456 0.0003 0.0003

N

- PropylCyc

- C6 4878.839 0.0006 4878.839 390 3.073 975.7652 975.7652 0.0087 975.7564 975.7565 0.0087 0.0087

1

- 2

- 3

- Mesitylene 641.6508 0 641.6508 128.3304 513.3206 513.3206 0.0191 513.3015 513.3015 0.0191 0.0191

N

- ButylCycHexane 233.9575 0 233.9575 125.987 107.9707 107.9707 0.0016 107.9691 107.9691 0.0016 0.0016

1

- 2

- DiC1

- 3C2

- Bz 1592.733 0 1592.733 135.2346 1457.499 1457.499 0.0871 1457.412 1457.412 0.0871 0.0871

Cis

- Decalin 3184.483 0 3184.484 1024.104 2160.379 2160.379 0.0983 2160.281 2160.281 0.0983 0.0984

Dimethyl

- C11 8633.395 0 8633.397 1406.6 76 7226.72 7226.72 0.0961 7226.625 7226.625 0.0961 0.0961

1

- 2

- 4

- triethylbe 3186.105 0 3186.106 49.8727 3136.233 3136.233 0.2496 3135.983 3135.984 0.2496 0.2497

1

- 1

- Bicyclohexyl 238.9565 0 238.9565 4.0338 234.9227 234.9227 0.035 234.8877 234.8877 0.035 0. 035

Diphenyl 4383.133 0 4383.134 2.1702 4380.966 4380.966 6.4693 4374.497 4374.496 6.4693 6.4695

diamantane 8633.229 0 8633.231 174.8519 8458.38 8458.38 15.8027 8442.577 8442.577 15.8027 15.8052

Phenanthrene 6087.884 0 6087.887 0 6087.889 6087.889 2653. 667 3434.22 3434.221 2653.668 2653.667

N

- C15

- CycloC5 291.442 0 291.442 0 291.4421 291.4421 3.1626 288.2794 288.2794 3.1626 3.1626

Chrysene 6120.283 0 6120.284 0 6120.288 6120.288 6120.288 0.0001 0.0001 6120.288 6120.286

1

- Pentene 0 0 0 0 0 0 0 0 0 0 0

O

- Xylene 0 0 0 0 0 0 0 0 0 0 0

Cyclopentane 0 0 0 0 0 0 0 0 0 0 0.0552

N

- Pentadecane 0 0 0 0 0 0 0 0 0 0 0

N

- Octadecane 0 0 0 0 0 0 0 0 0 0 0

Tetralin 0 0 0 0 0 0 0 0 0 0 0.0439

Ethylbenzene 0 0 0 0 0 0 0 0 0 0 0

Benzene 0 0 0 0 0 0 0 0 0 0 0

N

- Buta ne 0 0 0 0 0 0 0 0 0 0 28.8114

Toluene 0 0 0 0 0 0 0 0 0 0 0

1

- 3

- 5

- Mesitylene 0 0 0 0 0 0 0 0 0 0 0

P

- Xylene 1732.494 0.3188 1732.176 1248.131 484.0454 484.0454 0.0088 484.0365 484.0366 0.0088 0.0088

N

- Propylbenzene 0.0102 0 0.0102 0.0064 0.0038 0.0038 0 0.0038 0.0038 0 0

N

- Butylbenzene 0 0 0 0 0 0 0 0 0 0 0

O

- Ethyltoluene 0 0 0 0 0 0 0 0 0 0 0

C9H18 0 0 0 0 0 0 0 0 0 0 13.652

MthCyclohexane 0 0 0 0 0 0 0 0 0 0 38.0712

# A.11 Table A.3 . (contd)

Stream No. 510 511 515 523 524 526 529 531 532 536 540

T emp F 1247.858 402.8347 110 110 118.5025 110 110 109.9931 77 95 109.9931

Pres psig 1275 1275 1255 1255 1300 1255 1255 1 300 1300 1

Enth MMBtu/h 5.6979

- 4.3021

- 7.9143

- 0.09247

- 0.06215

- 0.01709

- 7.8047

- 0.00062

- 0.15399

- 0.08641

- 7.8041

Vapor mass fr action 1 0.6392 0.11945 1 1 1 0 1 1 1 0

Total lb/h 11010.12 11010.12 11010.12 1110 1110 205.1373 9694.978 0.9811 1100 1100 9694

Flowrates in lb/h

Water 0 0 0 0 0 0 0 0 0 0 0

Hybrid Poplar Ch 0 0 0 0 0 0 0 0 0 0 0

Hydrogen 1219.721 1219.721 1219.77 1029.363 1029.363 190.2349 0.172 0.1702 1100 1100 0.0018

Carbon Dioxide 0 0 0 0 0 0 0 0 0 0 0

Methane 0 0 0 0 0 0 0 0 0 0 0

Ethane 0 0 0 0 0 0 0 0 0 0 0

Propane 0 0 0 0 0 0 0 0 0 0 0

I

- Butane 0 0 0 0 0 0 0 0 0 0 0

2

- 5

- Xylenol 0.0083 0.0083 0. 0083 0.0001 0.0001 0 0.0083 0 0 0 0.0083

N

- Heptane 0 0 0 0 0 0 0 0 0 0 0

1

- ts

- 35

- 3C1CycC6 0.0016 0.0016 0.0016 0 0 0 0.0016 0 0 0 0.0016

3

- 3

- 5

- TriMth

- C7 0.0003 0.0003 0.0003 0 0 0 0.0003 0 0 0 0.0003

N

- PropylCyc

- C6 0.0087 0.0087 0.0087 0 0 0 0.0087 0 0 0 0.0087

1

- 2

- 3

- Mesitylene 0.0191 0.0191 0.0191 0 0 0 0.0191 0 0 0 0.0191

N

- ButylCycHexane 0.0016 0.0016 0.0016 0 0 0 0.0016 0 0 0 0.0016

1

- 2

- DiC1

- 3C2

- Bz 0.0871 0.0871 0.0871 0 0 0 0.0871 0 0 0 0.0871

Cis

- Decalin 0.0984 0.0984 0.0984 0 0 0 0.0983 0 0 0 0.0983

Dimethyl

- C11 0.0961 0.0961 0.0961 0 0 0 0.0961 0 0 0 0.0961

1

- 2

- 4

- triethylbe 0.2497 0.2497 0.2497 0 0 0 0.2496 0 0 0 0.2496

1

- 1

- Bicyclohexyl 0.035 0.035 0.035 0 0 0 0.035 0 0 0 0.035

Diphenyl 6.4695 6.4695 6.4695 0.0002 0.0002 0 6.4693 0 0 0 6. 4693

diamantane 15.8052 15.8052 15.8052 0.0026 0.0026 0.0005 15.8022 0 0 0 15.8022

Phenanthrene 0 0 0 0 0 0 0 0 0 0 0

N

- C15

- CycloC5 0.0032 0.0032 0.0032 0 0 0 0.0032 0 0 0 0.0032

Chrysene 0 0 0 0 0 0 0 0 0 0 0

1

- Pentene 0 0 0 0 0 0 0 0 0 0 0

O

- Xylene 0 0 0 0 0 0 0 0 0 0 0

Cyclopentane 0.8451 0.8451 0.845 0.0552 0.0552 0.0102 0.7796 0.0006 0 0 0.779

N

- Pentadecane 2.3923 2.3923 2.3923 0 0 0 2.3923 0 0 0 2.3923

N

- Octadecane 0 0 0 0 0 0 0 0 0 0 0

Tetralin 196.8838 196.8838 196.8838 0.0439 0.0439 0.008 1 196.8318 0.0004 0 0 196.8314

Ethylbenzene 0 0 0 0 0 0 0 0 0 0 0

Benzene 0 0 0 0 0 0 0 0 0 0 0

N

- Butane 115.3507 115.3507 115.3013 28.8121 28.8121 5.3247 81.1645 0.2964 0 0 80.8681

Toluene 0 0 0 0 0 0 0 0 0 0 0

1

- 3

- 5

- Mesitylene 0 0 0 0 0 0 0 0 0 0 0

P

- Xylene 0.0088 0.0088 0.0088 0 0 0 0.0088 0 0 0 0.0088

N

- Propylbenzene 0 0 0 0 0 0 0 0 0 0 0

N

- Butylbenzene 0 0 0 0 0 0 0 0 0 0 0

O

- Ethyltoluene 0 0 0 0 0 0 0 0 0 0 0

C9H18 6782.503 6782.503 6782.505 13.652 13.652 2.523 6766.33 0.1299 0 0 6766.2

Mth Cyclohexane 2669.528 2669.528 2669.527 38.0711 38.0711 7.0359 2624.419 0.3836 0 0 2624.036

# A.12 Table A.3 . (contd)

Stream No. 544 547 601 602 735 737 1850 1851

Temp F 150 150 85.0536 85.0536 700.0002 496.55 90 105

Pres psig 1 15 35 35 645 645 45. 3041 40.3041

Enth MMBtu/h

- 21.85

- 15.53

- 39.134

- 10.149

- 28.903

- 33.135

- 129.5

- 129.21

Vapor mass fraction 0 0 0.98769 0.98769 1 1.00E

- 06 0 0

Total lb/h 28555.25 38367.25 15824.14 4103.631 5193.402 5193.402 19029.97 19029.96

Flowrates in l b/h

Water 0 0 249.8139 64.7836 5193.402 5193.402 19029.97 19029.96

Hybrid Poplar Ch 0 0 0 0 0 0 0 0

Hydrogen 0.0018 0 457.8133 118.7234 0 0 0 0

Carbon Dioxide 0 0 6947.662 1801.718 0 0 0 0

Methane 0 0 2656.017 688.7776 0 0 0 0

Ethane 0 0 1593.595 413.2626 0 0 0 0

Propane 0 0 1340.26 347.5659 0 0 0 0

I

- Butane 0 0 1189.417 308.4483 0 0 0 0

2

- 5

- Xylenol 3162.44 4585.968 1.8125 0.47 0 0 0 0

N

- Heptane 1596.764 0 1332.604 345.5806 0 0 0 0

1

- ts

- 35

- 3C1CycC6 4541.51 8 316.1217 27.7666 7.2006 0 0 0 0

3

- 3

- 5

- TriMth

- C7 1856.981 130.4455 5.7693 1.4961 0 0 0 0

N

- PropylCyc

- C6 3903.079 975.7582 10.933 2.8352 0 0 0 0

1

- 2

- 3

- Mesitylene 128.3341 513.3166 0.2678 0.0695 0 0 0 0

N

- ButylCycHexane 125.9872 107.9704 0.1 93 0.05 0 0 0 0

1

- 2

- DiC1

- 3C2

- Bz 135.2398 1457.493 0.3383 0.0877 0 0 0 0

Cis

- Decalin 1024.112 2160.371 1.5088 0.3913 0 0 0 0

Dimethyl

- C11 1406.674 7226.722 1.5272 0.3961 0 0 0 0

1

- 2

- 4

- triethylbe 49.8727 3136.233 0.253 0.0656 0 0 0 0

1

- 1

- Bicyclohexyl 4.0338 234.9227 0.0204 0.0053 0 0 0 0

Diphenyl 2.1701 4380.964 0.0869 0.0225 0 0 0 0

diamantane 174.8532 8458.38 1.6581 0.43 0 0 0 0

Phenanthrene 0 3434.22 0.0042 0.0011 0 0 0 0

N

- C15

- CycloC5 0 288.2826 0.0001 0 0 0 0 0

Chr ysene 0 0.0001 0 0 0 0 0 0

1

- Pentene 0 0 0 0 0 0 0 0

O

- Xylene 0 0 0 0 0 0 0 0

Cyclopentane 0.779 0 0.0005 0.0001 0 0 0 0

N

- Pentadecane 0 2.3923 0 0 0 0 0 0

N

- Octadecane 0 0 0 0 0 0 0 0

Tetralin 4.8831 191.9483 0.0003 0.0001 0 0 0 0

Ethylbenzene 0 0 0 0 0 0 0 0

Benzene 0 0 0 0 0 0 0 0

N

- Butane 80.868 0 0.2354 0.061 0 0 0 0

Toluene 0 0 0 0 0 0 0 0

1

- 3

- 5

- Mesitylene 0 0 0 0 0 0 0 0

P

- Xylene 1248.139 484.0367 4.1775 1.0833 0 0 0 0

N

- Propylbenzene 0.0064 0.0038 0 0 0 0 0 0

N

- Butylbenzene 0 0 0 0 0 0 0 0

O

- Ethyltoluene 0 0 0 0 0 0 0 0

C9H18 6484.484 281.7122 0.1032 0.0268 0 0 0 0

MthCyclohexane 2624.033 0 0.3046 0.079 0 0 0 0

# A.13 Figure A.4 . Flow Diagram for Hydrogen Generation by Steam Reforming

# A.14 Table A.4 . Hydrogen Generation by Steam Reforming Heat and Material Balance

Stream No. 250 532 500 501 502 503 505 5 07 508 509 512

Temp F 77 77 60 61.8717 700 700 1050 1562 549.1881 667.5483 150

Pres psig 300 300 400 400 399.5 399.5 389.5 359.5 354.5 344.5 329.5

Enth MMBtu/h

- 0.90996

- 0.15399

- 34.902

- 34.892

- 26.96

- 26.96

- 498.35

- 367.68

- 440.7

- 440.7

- 526.32

Vap or mass fraction 1 1 1 1 1 1 1 1 1 1 0.53835

Total lb/h 6500 1100 17196.06 17401.2 17401.2 17401.2 110098.3 110098.4 110098.4 110098.8 110098.8

Flowrates in lb/h

Water 0 0 0 0 0 0 88044.1 60993.9 60993.9 51908.3 51908.3

Hydrogen 6500.0 1100. 0 0.0 190.2 190.2 190.2 327.6 7450.3 7450.3 8466.9 8466.9

Carbon 0 0 0 0 0 0 0 0 0 0 0

Carbon Monoxide 0 0 0 0 0 0 0 17026.3 17026.3 2899.8 2899.8

Carbon Dioxide 0 0 0 0 0 0 2084.2 21749.5 21749.5 43945.3 43945.3

Methane 0 0 17196.1 17196.1 17196.1 171 96.1 17992.9 2878.1 2878.1 2878.1 2878.1

Ethylene 0 0 0 0 0 0 0 0 0 0 0

Ethane 0 0 0 0 0 0 478.0 0.1 0.1 0.1 0.1

Propylene 0 0 0 0 0 0 0 0 0 0 0

Propane 0 0 0 0 0 0 401.9 0 0 0 0

I

- Butane 0 0 0 0 0 0 356.5 0 0 0 0

Formaldehyde 0 0 0 0 0 0 0 0.2 0.2 0 .2 0.2

2

- 5

- Xylenol 0 0 0 0 0 0 0.1 0 0 0 0

N

- Heptane 0 0 0 0 0 0 385.5 0 0 0 0

1

- ts

- 35

- 3C1CycC6 0 0 0 0 0 0 6.9 0 0 0 0

3

- 3

- 5

- TriMth

- C7 0 0 0 0 0 0 1.3 0 0 0 0

N

- PropylCyc

- C6 0 0 0 0 0 0 2.6 0 0 0 0

1

- 2

- 3

- Mesitylene 0 0 0 0 0 0 0 0 0 0 0

N

- ButylCycH exane 0 0 0 0 0 0 0 0 0 0 0

1

- 2

- DiC1

- 3C2

- Bz 0 0 0 0 0 0 0 0 0 0 0

Cis

- Decalin 0 0 0 0 0 0 0.2 0 0 0 0

Dimethyl

- C11 0 0 0 0 0 0 0.1 0 0 0 0

1

- 2

- 4

- triethylbe 0 0 0 0 0 0 0 0 0 0 0

1

- 1

- Bicyclohexyl 0 0 0 0 0 0 0 0 0 0 0

Diphenyl 0 0 0 0 0 0 0 0 0 0 0

d iamantane 0 0 0 0 0 0 0.1 0 0 0 0

Phenanthrene 0 0 0 0 0 0 0 0 0 0 0

N

- C15

- CycloC5 0 0 0 0 0 0 0 0 0 0 0

Cyclopentane 0 0 0 0 0 0 0 0 0 0 0

N

- Pentadecane 0 0 0 0 0 0 0 0 0 0 0

N

- Octadecane 0 0 0 0 0 0 0 0 0 0 0

Tetralin 0 0 0 0 0 0 0 0 0 0 0

Ethylbe nzene 0 0 0 0 0 0 0 0 0 0 0

Benzene 0 0 0 0 0 0 0 0 0 0 0

N

- Butane 0 0 0 5.3 5.3 5.3 5.4 0 0 0 0

Toluene 0 0 0 0 0 0 0 0 0 0 0

1

- 3

- 5

- Mesitylene 0 0 0 0 0 0 0 0 0 0 0

P

- Xylene 0 0 0 0 0 0 1.1 0 0 0 0

C9H18 0 0 0 2.5 2.5 2.5 2.5 0 0 0 0

MthCyclohexane 0 0 0 7.0 7.0 7.0 7.1 0 0 0 0

# A.15 Table A.4 . (contd)

Stream No. 513 514 520 526 601 603 604 605 609 610 612

Temp F 150 150 150 110 85.0536 85.0536 85.0536 85.0536 256.2011 60 550

Pres psig 329.5 329.5 329.5 1255 35 35 35 35 400 0.001 2

Enth MMBtu/h

- 183.1

- 343.22 1.9013 0.009921

- 39.131

- 38.37

- 0.76068

- 26.859

- 11.172

- 0.97388 26.657

Vapor mass fraction 1 0 1 1 0.98944 1 0 1 1 1 1

Total lb/h 59271.69 50827.07 7620.208 205.1373 15824.14 15657.06 167.0814 10959.94 4697.12 230000 230000

Flowrates in lb/h

Water 1081.2 50827.1 0 0 249.8 147.0 102.8 102.9 44.1 0 0

Hydrogen 8466.9 0 7620.2 190.2 457.8 457.8 0.0 320.5 137.3 0 0

Carbon 0.0 0 0 0 0 0 0 0 0 0 0

Carbon Monoxide 2899.8 0 0 0 0 0 0 0 0 0 0

Carbon Dioxide 43945.3 0 0 0 6947.7 6947 .3 0.4 4863.1 2084.2 0 0

Methane 2878.1 0 0 0 2656.0 2656.0 0.0 1859.2 796.8 0 0

Ethylene 0 0 0 0 0 0 0 0 0 0 0

Ethane 0.1 0 0 0 1593.6 1593.4 0.1 1115.4 478.0 0 0

Propylene 0 0 0 0 0 0 0 0 0 0 0

Propane 0 0 0 0 1340.3 1339.8 0.4 937.9 401.9 0 0

I

- Bu tane 0 0 0 0 1189.4 1188.5 1.0 831.9 356.5 0 0

Formaldehyde 0.2 0 0 0 0 0 0 0 0 0 0

2

- 5

- Xylenol 0.0 0 0 0 1.8 0.4 1.4 0.3 0.1 0 0

N

- Heptane 0 0 0 0 1332.6 1285.0 47.6 899.5 385.5 0 0

1

- ts

- 35

- 3C1CycC6 0 0 0 0 27.8 23.1 4.7 16.1 6.9 0 0

3

- 3

- 5

- TriMth

- C7 0 0 0 0 5.8 4.3 1.5 3.0 1.3 0 0

N

- PropylCyc

- C6 0 0 0 0 10.9 8.7 2.3 6.1 2.6 0 0

1

- 2

- 3

- Mesitylene 0 0 0 0 0.3 0.2 0.1 0.1 0 0 0

N

- ButylCycHexane 0 0 0 0 0.2 0.1 0.1 0.1 0 0 0

1

- 2

- DiC1

- 3C2

- Bz 0 0 0 0 0.3 0.1 0.2 0.1 0 0 0

Cis

- Decalin 0 0 0 0 1.5 0.7 0.8 0.5 0.2 0 0

Dimethyl

- C11 0 0 0 0 1.5 0.2 1.3 0.1 0.1 0 0

1

- 2

- 4

- triethylbe 0 0 0 0 0.3 0 0.2 0 0 0 0

1

- 1

- Bicyclohexyl 0 0 0 0 0.0 0 0 0 0 0 0

Diphenyl 0 0 0 0 0.1 0 0.1 0 0 0 0

diamantane 0 0 0 0 1.7 0.3 1.4 0.2 0.1 0 0

Phenanthrene 0 0 0 0 0 0 0 0 0 0 0

N

- C15

- CycloC5 0 0 0 0 0 0 0 0 0 0 0

Cyclopentane 0 0 0 0 0 0 0 0 0 0 0

N

- Pentadecane 0 0 0 0 0 0 0 0 0 0 0

N

- Octadecane 0 0 0 0 0 0 0 0 0 0 0

Tetralin 0 0 0 0 0 0 0 0 0 0 0

Ethylbenzene 0 0 0 0 0 0 0 0 0 0 0

Benzene 0 0 0 0 0 0 0 0 0 0 0

N

- But ane 0 0 0 5.3 0.2 0.2 0.0 0.2 0.1 0 0

Toluene 0 0 0 0 0 0 0 0 0 0 0

1

- 3

- 5

- Mesitylene 0 0 0 0 0 0 0 0 0 0 0

P

- Xylene 0 0 0 0 4.2 3.6 0.6 2.5 1.1 0 0

C9H18 0 0 0 2.5 0.1 0.1 0 0.1 0 0 0

MthCyclohexane 0 0 0 7.0 0.3 0.3 0 0.2 0.1 0 0

# A.16 Table A.4 . (cont d)

Stream No. 615 630 706 707 721 725 730 733 734 735 736

Temp F 150 311.6723 276.468 276.468 497.3775 497.3795 497.3795 700.0002 700 700 700

Pres psig 329.5 1.5 670 670 650 650 650 645 645 645 645

Enth MMBtu/h

- 185.86

- 447.3

- 1133.3

- 0.00703

- 1371. 1

- 1316.6

- 675.88

- 121.47

- 486.09

- 28.903

- 22.468

Vapor mass fraction 0.98557 1 0 0 0 0.355 1 1 1 1 1

Total lb/h 51651.48 292778.2 171067 1.0604 214800.3 214800 119221.6 21991.57 88000 5193 4037

Flowrates in lb/h

Water 1081.2 28603.1 171067. 0 1.1 214800.3 214800.0 119221.6 21991.6 88000 5193.0 4037.0

Hydrogen 846.7 0 0 0 0 0 0 0 0 0 0

Carbon 0 0 0 0 0 0 0 0 0 0 0

Carbon Monoxide 2899.8 0 0 0 0 0 0 0 0 0 0

Carbon Dioxide 43945.3 78010.5 0 0 0 0 0 0 0 0 0

Methane 2878.1 0 0 0 0 0 0 0 0 0 0

Ethylene 0 0 0 0 0 0 0 0 0 0 0

Ethane 0.1 0 0 0 0 0 0 0 0 0 0

Propylene 0 0 0 0 0 0 0 0 0 0 0

Propane 0 0 0 0 0 0 0 0 0 0 0

I

- Butane 0 0 0 0 0 0 0 0 0 0 0

Formaldehyde 0 0 0 0 0 0 0 0 0 0 0

2

- 5

- Xylenol 0 0 0 0 0 0 0 0 0 0 0

N

- Heptane 0 0 0 0 0 0 0 0 0 0 0

1

- ts

- 35

- 3C1CycC6 0 0 0 0 0 0 0 0 0 0 0

3

- 3

- 5

- TriMth

- C7 0 0 0 0 0 0 0 0 0 0 0

N

- PropylCyc

- C6 0 0 0 0 0 0 0 0 0 0 0

1

- 2

- 3

- Mesitylene 0 0 0 0 0 0 0 0 0 0 0

N

- ButylCycHexane 0 0 0 0 0 0 0 0 0 0 0

1

- 2

- DiC1

- 3C2

- Bz 0 0 0 0 0 0 0 0 0 0 0

Cis

- Decali n 0 0 0 0 0 0 0 0 0 0 0

Dimethyl

- C11 0 0 0 0 0 0 0 0 0 0 0

1

- 2

- 4

- triethylbe 0 0 0 0 0 0 0 0 0 0 0

1

- 1

- Bicyclohexyl 0 0 0 0 0 0 0 0 0 0 0

Diphenyl 0 0 0 0 0 0 0 0 0 0 0

diamantane 0 0 0 0 0 0 0 0 0 0 0

Phenanthrene 0 0 0 0 0 0 0 0 0 0 0

N

- C15

- CycloC5 0 0 0 0 0 0 0 0 0 0 0

Cyclopentane 0 0 0 0 0 0 0 0 0 0 0

N

- Pentadecane 0 0 0 0 0 0 0 0 0 0 0

N

- Octadecane 0 0 0 0 0 0 0 0 0 0 0

Tetralin 0 0 0 0 0 0 0 0 0 0 0

Ethylbenzene 0 0 0 0 0 0 0 0 0 0 0

Benzene 0 0 0 0 0 0 0 0 0 0 0

N

- Butane 0 0 0 0 0 0 0 0 0 0 0

Toluene 0 0 0 0 0 0 0 0 0 0 0

1

- 3

- 5

- Mesitylene 0 0 0 0 0 0 0 0 0 0 0

P

- Xylene 0 0 0 0 0 0 0 0 0 0 0

C9H18 0 0 0 0 0 0 0 0 0 0 0

MthCyclohexane 0 0 0 0 0 0 0 0 0 0 0

# A.17 Table A.4 . (contd)

Stream No. 743 754

Temp F 499.8419 400

Pres psig 66 5 665

Enth MMBtu/h

- 501.74

- 536.23

Vapor mass fraction 1 0

Total lb/h 88500 82567.1

Flowrates in lb/h

Water 88500.0 82567.1

Hydrogen 0 0

Carbon 0 0

Carbon Monoxide 0 0

Carbon Dioxide 0 0

Methane 0 0

Ethylene 0 0

Ethane 0 0

Propylene 0 0

Pr opane 0 0

I

- Butane 0 0

Formaldehyde 0 0

2

- 5

- Xylenol 0 0

N

- Heptane 0 0

1

- ts

- 35

- 3C1CycC6 0 0

3

- 3

- 5

- TriMth

- C7 0 0

N

- PropylCyc

- C6 0 0

1

- 2

- 3

- Mesitylene 0 0

N

- ButylCycHexane 0 0

1

- 2

- DiC1

- 3C2

- Bz 0 0

Cis

- Decalin 0 0

Dimethyl

- C11 0 0

1

- 2

- 4

- triethylbe 0 0

1

- 1

- Bicyclohexyl 0 0

Diphenyl 0 0

diamantane 0 0

Phenanthrene 0 0

N

- C15

- CycloC5 0 0

Cyclopentane 0 0

N

- Pentadecane 0 0

N

- Octadecane 0 0

Tetralin 0 0

Ethylbenzene 0 0

Benzene 0 0

N

- Butane 0 0

Toluene 0 0

1

- 3

- 5

- Mesitylene 0 0

P

- Xylene 0 0

C9H18 0 0

MthCyclohexane 0 0

# A.18 Figure A.5 . Flow Diagram for the Steam Reformer Steam Drum

# A.19 Table A.5 . Steam Reformer Steam Drum Heat and Material Balance

Stream No. 707 720 721 725 730 743

Temp F 276.5 497.4 497.4 497.4 497.4 499.8

Pres psig 670 650 650 650 650 665

Enth MMBtu/h

- 0.007

- 1371.1

- 1371.1

- 1316.6

- 675.88

- 501.74

Vapor mass

fraction 0 0 0 0.355 1 1

Total lb/h 1.1 214799.2 214800.3 21 4800.0 119221.6 88500.0

Flowrates in lb/h

Water 1.1 214799.2 214800.3 214800.0 119221.6 88500.0

Stream No. 754 760 762 765 781 790

Temp F 400.0 497.4 497.4 497.4 497.4 497.4

Pres psig 665 650 650 650 650 650

Enth MMBtu/h

- 536 .23

- 91.839

- 21.543

- 23.81

- 47.62

- 122.87

Vapor mass

fraction 0 1 1 1 1 0

Total lb/h 82567.1 16200.0 3800.0 4200.0 8400.0 19250.0

Flowrates in lb/h

Water 82567.1 16200.0 3800.0 4200.0 8400.0 19250.0

# A.20 Figure A.6 . Flow Diagram for Condensate Collection

# A.21 Table A.6 . Condensate Collection Heat and Material Balance

Stream No. 514 700 706 707 737 738

Temp F 150.0 60.0 276.5 276.5 496.6 496.6

Pr es psig 329.5 20 670 670 645 645

Enth MMBtu/h

- 343.22

- 386.06

- 1133.3

- 0.0070253

- 33.135

- 25.758

Vapor mass

fraction 0 0 0 0 1.00E

- 06 1.00E

- 06

Total lb/h 50827.1 56419.0 171067.0 1.1 5193.4 4037.1

Flowrates in lb/h

Water 50827.1 56419.0 171067 .0 1.1 5193.4 4037.1

Stream No. 761 763 766 768 781

Temp F 495.6 497.4 496.6 366.0 497.4

Pres psig 645 650 645 150 650

Enth MMBtu/h

- 103.44

- 24.193

- 26.71

- 143.64

- 47.62

Vapor mass

fraction 0 0.023151 0.034147 0 1

Total lb/h 1 6200.0 3800.0 4200.0 21992.0 8400.0

Flowrates in lb/h

Water 16200.0 3800.0 4200.0 21992.0 8400.0

# Appendix B

Equipment Cost Details

# B.1 

# Appendix B

Equipment Cost Details

Table B.1 lists the equipment cost details. Table B.2 lists the specifications for the standard equipment.B.2 

# Table B.1 . Equipment Capital Cost Detail

#

Reqd Equipment Name Original

Stream

Metric New

Stream

Metric Scaling

U nits Size

Ratio Original

Equip Cost

(per unit) Base Year Scaling

Exp I nstall

Factor Scaled and

Installed Cost

in Base Year Installed Cost

in 2007$ Bare Equip

cost in 2007$ Equip

Cost

Source

A100 Pyrolysis Oil

1 Feedstock Handling 500 2,000 tpd 4.00 $5,570,000 2003 0.7 2.47 14,699,318 19,212,106 $7,778,181 1

1 CFB Pyrolyzer 500 2,000 tpd 4.00 $3,392,000 2003 0.7 2.47 8,951,542 11,699,724 $4,736,731 1

1 Quench 500 2,000 tpd 4.00 $1,940,000 2003 0.7 2.47 5,119,691 6,691,470 $2,709,097 1

1 Heat recovery 500 0 tpd 0.00 $1,140,000 2003 0.7 2.47 0 0 $0 1

1 Produ ct recovery and storage 500 2,000 tpd 4.00 $800,000 2003 0.7 2.47 2,111,213 2,759,369 $1,117,153 1

1 Recycle 500 2,000 tpd 4.00 $1,380,000 2003 0.7 2.47 3,641,842 4,759,911 $1,927,090 1

1 Steam and Power production 500 0 tpd 0.00 $3,160,000 2003 0.7 2 .47 0 0 $0 1

1 Utilities 500 0 tpd 0.00 $3,130,000 2003 0.7 2.47 0 0 $0 1

$34,523,605 45,122,581 $18,268,251

Equipment Contingency 35% $15,792,903 $6,393,888

60,915,484 $24,662,139

Cost sources 1) Ringer et al. 2006

A 200 Pyrolysis Oil Upgrading

2 Feed Booster Pump 1,000 2,000 tpd 2.00 $30,000 2004 0.65 2.47 232,551 275,070 $111,364 2

2 Feed Pump 1,000 2,000 tpd 2.00 $122,000 2004 0.65 2.47 945,706 1,118,617 $452,882 2

1 1 st 

### stage React or 2,038 1,959 liq scfh 0.96 $860,700 1Q 2007 0.65 2.47 2,125,929 2,120,904 $858,666 3

1 2 nd 

### stage Reactor 2,038 1,959 liq scfh 0.96 $9,018,000 1Q 2007 0.65 2.47 19,852,131 22,221,810 $8,996,684 3

1 2 nd 

### stage Three Phase

Separator 174,497 161,375 lb/h 0. 92 $673,300 1Q 2005 0.65 2.47 1,663,113 1,777,236 $719,529 3

1 1 st 

### Feed/Product Heat

Exchanger 59.85 38.38 mmbtuh 0.64 $647,800 1Q 2005 0.65 2.47 1,628,026 1,347,791 $545,664 3

1 2 nd 

### Feed/Product Heat

Exchanger 33.30 1.22 mmbtuh 0.04 $622,200 1Q 2005 0.6 5 2.47 1,537,134 201,410 $81,542 3

1 Fired Heater 8.35 40.00 mmbtuh 8.18 $378,000 2004 0.65 2.47 933,660 3,057,432 $1,237,827 2

1 Air Cooler 40.38 28.10 mmbtuh 0.70 $228,000 1Q 2005 0.65 2.47 548,093 500,254 $202,532 3

1 Product Trim Cooler 8.35 8.35 mm btuh 1.00 $128,700 1Q 2005 0.65 2.47 317,889 357,424 $144,706 3

1 PSA 10 15 mmscfd

H2 1.50 $1,750,000 2004 0.65 2.47 4,838,004 2,694,155 $1,090,751 2

1 Product Flash Drum 71,462 69,520 lb/h 0.97 $38,800 1Q 2005 0.65 2.47 95,836 105,842 $42,851 3

1 Produ ct Pump 1,000 2,000 tpd 2.00 $39,000 2004 0.65 2.47 151,158 357,591 $144,774 2

1 Hydrogen Compressor 2,840 2,840 acfm 1.00 $3,869,400 1Q 2005 0.65 2.47 9,557,418 10,746,045 $4,350,625 3

Subtotal $43,497,310 $46,881,581 $18,980,397

Equipment Co ntingency 15% $7,032,237 $2,847,060

$53,913,818 $21,827,457

Cost sources 2) Vendor quote, 2004; 3) Aspen IcarusB.3 

# Table B.1 . (contd)

#

Reqd Equipment Name Original

Stream

Metric New

Stream

Metric Scaling

units Size

Ratio Original

Eq uip Cost

(per unit) Base Year Scaling

Exp Install

Factor Scaled and

Installed Cost

in Base Year Installed Cost

in 2007$ Bare Equip cost

in 2007$ Equip

Cost

Source

A300/500 Hydrocra c king and

Separations

1 T

- 305 Debutanizer 69,495 69,495 lb/h 1 .00 $55,100 1Q 2007 0.65 2.47 217,113 222,238 $89,975 3

1 T

- 310 Naphtha Splitter 70,500 66,028 lb/h 0.94 $183,700 1Q 2005 0.65 2.47 434,817 488,894 $197,933 3

1 T

- 315 Diesel Splitter 45,548 46,668 lb/h 1.02 $65,500 1Q 2005 0.65 2.47 164,360 184,801 $7 4,818 3

1 T

- 540 9,680 9,694 lb/h 1.00 $123,800 1Q 2005 0.65 2.47 306,073 344,139 $139,327 3

1 T

- 305 Debutanizer Reboiler 11.60 11.60 mmbtuh 1.00 $48,600 1Q 2007 0.65 2.47 120,042 122,876 $49,747 3

1 T

- 310 Reboiler 2.90 2.90 mmbtuh 1.00 $27,200 1Q 20 07 0.65 2.47 67,184 68,770 $27,842 3

1 T

- 315 Reboiler 4.77 10.20 mmbtuh 2.14 $24,900 1Q 2005 0.65 2.47 100,797 113,333 $45,884 3

1 T

- 540 Reboiler 2.63 2.63 mmbtuh 1.00 $22,400 1Q 2007 0.65 2.47 55,328 56,634 $22,929 3

1 T

- 305 Debutanizer Condenser 1.4 8 1.48 mmbtuh 1.00 $33,200 1Q 2007 0.65 2.47 82,004 83,940 $33,984 3

1 T

- 310 Condenser 5.62 5.62 mmbtuh 1.00 $43,200 1Q 2007 0.65 2.47 106,704 109,223 $44,220 3

1 T

- 315 Condenser 13.40 13.40 mmbtuh 1.00 $55,800 1Q 2007 0.65 2.47 137,826 141,079 $57,11 7 3

1 T

- 540 Condenser 2.65 2.65 mmbtuh 1.00 $79,300 1Q 2007 0.65 2.47 195,871 200,495 $81,172 3

1 T

- 305 Reflux Drum 3655 3466 lb/h 0.95 $7,000 1Q 2005 0.65 2.47 16,703 18,781 $7,604 3

1 T

- 310 Reflux Drum 55 49 gpm 0.88 $17,500 1Q 2005 0.65 2.47 39,83 2 44,786 $18,132 3

1 T

- 315 Reflux Drum 82 89 gpm 1.08 $16,500 1Q 2005 0.65 2.47 42,890 48,224 $19,524 3

1 T

- 540 Reflux Drum 24 23 gpm 0.96 $11,500 1Q 2005 0.65 2.47 27,630 31,066 $12,577 3

1 Debutanizer Feed Preheater 6.66 3.29 mmbtuh 0.49 $29,000 1Q 2005 0.65 2.47 45,291 50,924 $20,617 3

1 T315 Feed Preheater 6.66 4.23 mmbtuh 0.64 $29,000 1Q 2005 0.65 2.47 53,328 59,961 $24,276 3

1 Hydrocracker Unit : 2,250 590 bpd fd 0.26 $30,000,000 1Q 2005 0.65 2.47 12,567,718 14,130,726 $5,720,942 4

H5 08 Fired heater Included in hydrocracker unit cost

R510 Hydrocracker Vessel Included in hydrocracker unit cost

E513 Feed/product exchanger Included in hydrocracker unit cost

E516 air cooler Included in hydrocracker u nit cost

E517 trim cooler Included in hydrocracker unit cost

V520 HP flash Included in hydrocracker unit cost

V530 LP flash Included in hydrocracker unit cost

1 Naphtha Product Cooler 1.4 1.45 mmbtuh 1.04 $46, 300 1Q 2005 0.65 2.47 116,999 131,550 $53,259 3

1 Diesel Product Cooler 3.1 4.93 mmbtuh 1.59 $44,600 1Q 2005 0.65 2.47 148,935 167,458 $67,797 3

16,819,896 6,809,675

Subtotal $2,522,984 $1,021,451

15% $19,342,881 $7,831,1 26

Cost sources 3) Aspen Icarus; 4) Marker et al. 2005B.4 

# Table B.1 . (contd)

#

Reqd Equipment Name Original

Stream

Metric New

Stream

Metric Scaling

units Size

Ratio Original

Equip Cost

(per unit) Base

Year Scaling

Exp Install

Factor Scal ed and

Installed Cost

in Base Year Installed Cost

in 2007$ Bare Equip cost

in 2007$ Equip

Cost

Source

A 700 Utilities and Auxiliaries

1 Wastewater treatment incinerator 1,000 0 tpd 0.00 $442,000 2004 1,000 2.47 0 0 $0 5

1 WWT anaerobic/aerob ic digestion 1,000 2,000 tpd 2.00 $1,554,000 2004 1,000 2.47 2,438,487 2,884,336 $1,167,747 5

1 Wastewater storage 593,849 672,840 gallons 1.13 226,710 1Q 2005 593,849 607,324 682,855 $276,459

1 Field

- Erected CTW w/pumps, etc 10,400 7,54 9 gpm 0.73 $352,000 1Q 2005 0.78 2.47 274,164 308,261 $124,802 3

1 Plant Air Compressor 2,000 2,000 tpd 1.00 $32,376 2002 0.34 2.47 79,969 106,212 $43,001 6

1 Hydraulic Truck Dump with Scale 2,000 2,000 tpd 1.00 $80,000 1998 0.6 2.47 197 ,600 266,553 $107,916 6

1 Firewater Pump 2,000 2,000 tpd 1.00 $18,400 1997 0.79 2.47 45,448 61,783 $25,013 6

1 Diesel Pump 2,000 2,000 tpd 1.00 $6,100 1997 0.79 2.47 15,067 20,482 $8,292 6

1 Instrument Air Dryer 2,000 2,000 tpd 1.00 $8,349 2002 0.6 2 .47 20,622 27,389 $11,089 6

1 Plant Air Receiver 2,000 2,000 tpd 1.00 $7,003 2002 0.72 2.47 17,297 22,974 $9,301 6

1 Firewater Storage Tank 2,000 2,000 tpd 1.00 $166,100 1997 0.51 2.47 410,267 557,726 $225,800 6

1 Ammonia Pump included in Stream Refor mer Cost

1 Hydrazine Pump included in Stream Reformer Cost

1 Ammonia Storage Tank included in Stream Reformer Cost

1 Hydrazine Storage Tank included in Stream Reformer Cost

1 Flare included in Stream Reformer Cost

1 Feed Storage 1,056,846 1,056,846 gallons 1.00 470,000 1Q 2005 0.65 2.47 1,160,900 1,305,278 $528,452 3

1 Product Storage 558,000 681,114 gallons 1.22 320,384 1Q 2005 0.65 2.47 836,744 940,807 $380,894 3

1 Product Storage 558,000 712,728 gallons 1.28 320,384 1Q 2005 0.65 2.47 930,580 1,046,313 $423,609 3

Subtotal $8,230,968 $3,332,376

Equipment Contingency 15% $1,234,645 $499,856

$9,465,613 $3,832,232

Cost sources 3) Aspen Icarus; 5) Beckman et al. 1990; 6) Phillips et al. 2007B.5 

# Table B.1 . (contd)

#

Reqd Equipment Name Original

Stream

Metric New

Stream

Metric Scaling

units Size

Ratio Original

Equip Cost

(per unit) Base

Year Scaling

Exp Install

Factor Scaled and

Installed Cost

in Base Year Total Project

Investment Equip

Cost

Source

A 600 Hydrogen Plant

1 Steam Reformer System w/ 24.5 34.4 mmscfd

H2 1.40 $69,900,000 May

2007 0.65 2.47 87,152,973 86,106,854 7

associated OSBL

Subtotal 87,152,973 86,106,854

Eq uipment Contingency 0% 0 0

87,152,973 86,106,854

Cost sources 7 ) SRI 2007

# B.6 Table B.2 . Standard Equipment Specifications

Tag Description T, °F P, psig Type and Size Metall urgy

R

- 205 1 st 

# stage Reactor 750 2530 Dia meter , ft

Len gth , ft 8

41 316 SS clad

R

- 215 2 nd 

# stage Reactor 780 2530 Dia meter , ft

Len gth , ft 16

80 316 SS clad

V

- 222 Three

- Phase Separator 100 700 Dia meter , ft

Len gth , ft 10

58 316 SS

E

- 202 1 st 

# Feed/Product Heat

Exchanger 750 2530 Shell & tube

floating head Area, ft 2



# Duty, MMBtu/h 6712

59.9 316 SS

E

- 206 2 nd 

# Feed/Product Heat

Exchanger 750 2530 Shell & tube

floating head Area, ft2

Duty, MMBtu/h 6518

33.3 316 SS

H

- 210 Fired Heater 850 2530 Duty, MMBtu/h 8 .18 316 SS

E

- 220 Air Cooler 400 750 Air Fin Area, ft2

Duty, MMBtu/h 6170

40.4 316 SS

E

- 221 Product Trim Cooler 150 700 Shell & tube

floating head Area, ft2

Duty, MMBtu/h 3589

8.35 316 SS

V

- 225 Product Flash Drum 100 100 Dia meter , ft

Len gth , ft 7

21 316 SS

C

- 250 Hydrogen Compressor 300 inlet Reciprocating

motor drive 3 stage 2480 acfm

T

- 305 Debutanizer 50 440 Dia meter , ft

# trays 5

1 8 CS

T

- 310 Naphtha Splitter 50 500 Dia meter , ft

# trays 3

28 CS

T

- 315 Diesel Splitter 50 700 Dia meter , ft

# trays 5

8 CS

T

- 540 T

- 540 50 400 Dia meter , ft

# trays 3

28 CS

T

- 305 Debutanizer Reboiler 150 750 Shell & tube

floating head Area, ft 2



# Duty, MMBtu/h 975

11.6 Tubes 316/ CS Shell

E

- 310R Reboiler 150 700 Shell & tube

floating head Area, ft 2



# Duty, MMBtu/h 224

2 .9 Tubes 316/ CS Shell

E

- 315R Reboiler 150 500 Shell & tube

floating head Area, ft 2



# Duty, MMBtu/h 218

4.7 Tubes 316/ CS Shell

E

- 540R Reboiler 150 880 Shell & tube

floating head Area, ft 2



# Duty, MMBtu/h 127

2.63 Tubes 316/ CS Shell

# B.7 Table B.2 . (contd)

E

- 3 05C Debutanizer Condenser 150 200 Air Fin Area, ft2

Duty, MMBtu/h 209

1.5 CS Tubes

E

- 310C Naphtha Splitter Condenser 150 200 Air Fin Area, ft2

Duty, MMBtu/h 503

5.6 CS Tubes

E

- 315C Diesel Splitter Condenser 150 300 Air Fin Area, ft2

Duty, MMBtu/h 670

13 .4 CS Tubes

E

- 540C T

- 540 Condenser 150 400 Air Fin Area, ft2

Duty, MMBtu/h 1680

2.7 CS Tubes

D

- 305 Debutanizer Reflux Drum 15 250 Horizontal, 5

min holdup Dia meter , ft

Len gth , ft 2

6 CS

D

- 310 Naphtha Splitter Reflux Drum 15 350 Horizontal, 5

min holdup Dia meter , ft

Len gth , ft 4.8

14.3 CS

D

- 315 Diesel Splitter Reflux Drum 15 425 Horizontal, 5

min holdup Dia meter , ft

Len gth , ft 4.5

13.5 CS

D

- 540 T

- 540 Reflux Drum 15 250 Horizontal, 5

min holdup Dia meter , ft

Len gth , ft 3.4

10.3 CS

E

- 305 Debutanizer Feed Preheater 150 530 Pre

- engineered Area, ft 2



# Duty, MMBtu/h 83

3.33 CS

E

- 311 Diesel Splitter Feed Preheater 150 Shell & tube

floating head Area, ft2

Duty, MMBtu/h 574

3.55 CS

E

- 542 Naphtha Product Cooler 50 475 Air Fin Area, ft 2



# Duty, MMBtu/h 838 CS

E

- 5 43 Diesel Product Cooler 50 475 Air Fin Area, ft 2



# Duty, MMBtu/h 670

3.1 CS

# y

Pacific Northwest

NATIONAL LABORATORY

902 Battelle Boulevard

P.O. Box 999

Richland, WA 99352

1-888-375-PNNL (7665)

www.pnl.gov U.S

. DEPARTMENT OF 

# ENERGY