Solar Thermal Conversion of
Biomass to Methanol
Instructor: A. W. Weimer
CHEN 4520
Bernard Britt
Robert McGugan
Sarah Stoeck
Andrew Weidner
Vanessa Witte
1
Table of Contents
I. Executive Summary ...................................................................................................................... 7
II. Introduction ................................................................................................................................ 8
Project Description and Scope .................................................................................................... 8
Mission Statement ................................................................................................................... 8
Process Description ................................................................................................................. 8
III. Background Information .......................................................................................................... 11
Market for Renewable Methanol .............................................................................................. 11
Renewable Methanol from Biomass ......................................................................................... 13
Solar-Thermal Processing .......................................................................................................... 14
IV. Environmental, Health & Safety .............................................................................................. 16
Chemical Hazards ...................................................................................................................... 16
Health and Safety Considerations ............................................................................................. 20
Operator Safety ..................................................................................................................... 20
Licensure and Permits ........................................................................................................... 22
Environmental Considerations .................................................................................................. 24
Worst-Case Scenarios................................................................................................................ 29
Life Cycle Analysis...................................................................................................................... 32
Goal and Scope Definition ..................................................................................................... 32
Inventory Analysis.................................................................................................................. 34
Impact Assessment and Interpretation ................................................................................. 37
V. Project Premises ....................................................................................................................... 38
Design ........................................................................................................................................ 38
Biomass Pre-Processing ......................................................................................................... 38
Biomass Gasification and Methanol Production ................................................................... 39
Amine Scrubbing .................................................................................................................... 39
Methanol Purification ............................................................................................................ 39
Economics.................................................................................................................................. 40
VI. Approach.................................................................................................................................. 42
Hand Calculations ...................................................................................................................... 42
Heat of Reactions: Cellulose and Steam ................................................................................ 42
2
Heat of Reaction: Lignin and Steam ...................................................................................... 43
Waste Biomass Feed Estimation ........................................................................................... 43
Theoretical Energy Requirement for Solar-Thermal Reactor ................................................ 45
VII. Process Flow Diagrams with Material & Energy Balances ..................................................... 46
Biomass Pre-Processing............................................................................................................. 46
Process Description and PFD ................................................................................................. 46
Material Balances .................................................................................................................. 48
Heat Duty ............................................................................................................................... 49
Biomass Gasification ................................................................................................................. 49
Process Description and PFD ................................................................................................. 49
Heat Duty ............................................................................................................................... 61
Amine Scrubbing ....................................................................................................................... 62
Process Description and PFD ................................................................................................. 62
Material Balance .................................................................................................................... 70
Heat Duty ............................................................................................................................... 71
Product Separation & Post-Processing ..................................................................................... 71
Process Description and PFD ................................................................................................. 71
Material Balances .................................................................................................................. 73
Heat Duty ............................................................................................................................... 74
VIII. Process Description & Equipment Specifications .................................................................. 75
Generalized Equipment Design ................................................................................................. 75
Pumps .................................................................................................................................... 75
Shell-and-Tube Heat Exchangers ........................................................................................... 78
Vapor-Liquid Separators ........................................................................................................ 82
Pressure Vessels and Towers................................................................................................. 90
Cyclones ................................................................................................................................. 94
Biomass Pre-Processing............................................................................................................. 98
P-1 / SR-101 Shredding .......................................................................................................... 98
P-2 / RDR-101 Rotary Drying ............................................................................................... 100
P-3 / GR-101 Grinding .......................................................................................................... 103
P-4 / HP-101 Hopper............................................................................................................ 105
3
Solar Field and Tower .............................................................................................................. 110
Biomass Gasification ............................................................................................................... 112
Solar Reactor........................................................................................................................ 112
Zinc-Oxide Reactor .............................................................................................................. 118
Methanol Reactor ................................................................................................................ 119
Spray Quench Tank .............................................................................................................. 122
Heat Exchangers .................................................................................................................. 124
Compressor .......................................................................................................................... 126
Pumps .................................................................................................................................. 128
Vapor-Liquid Separators ...................................................................................................... 130
ZN-SPLIT Cyclone ................................................................................................................. 134
Amine Scrubbing ..................................................................................................................... 134
Pressure Vessels, Separators, and Towers .......................................................................... 134
Heat Exchangers .................................................................................................................. 138
Pumps .................................................................................................................................. 143
Product Separation & Post-Processing ................................................................................... 144
Separator V-100 ................................................................................................................... 144
Distillation Column T-100 .................................................................................................... 145
IX. Utility Summary and Heat Integration ................................................................................... 150
Utility Summary ....................................................................................................................... 150
Heat Integration ...................................................................................................................... 152
X. Estimation of Capital Investment and Total Product Cost...................................................... 157
Capital Investment .................................................................................................................. 157
Equipment Cost Summary ................................................................................................... 157
Operating Costs ....................................................................................................................... 172
Metrics of Plant Operation .................................................................................................. 172
Variable Operating Costs ..................................................................................................... 172
Fixed Operating Costs .......................................................................................................... 172
XI. Profitability Analysis............................................................................................................... 174
Profitability .............................................................................................................................. 174
Base Case ............................................................................................................................. 174
4
Modified Base Case: Life Cycle Analysis with Carbon Credit ............................................... 179
Sensitivity Analysis .................................................................................................................. 181
XII. Conclusion ............................................................................................................................. 188
XIII. References ........................................................................................................................... 189
XIII. Appendix .............................................................................................................................. 196
Appendix I: Engineering Calculations ...................................................................................... 196
Appendix I-A: Approach Calculations .................................................................................. 196
Appendix I-B: Material and Energy Balances ...................................................................... 200
Appendix I-C: Split Fraction for AG-CLEAN .......................................................................... 205
Appendix I-D: Adiabatic Temperature Rise for Methanol Reactor ..................................... 206
Appendix II: Design and Costing Spreadsheets ....................................................................... 207
Appendix II-A: Solar Field Design ......................................................................................... 207
Appendix II-B: Cyclone Design ............................................................................................. 208
Appendix II-C: Aspen PLUS Separator Design ...................................................................... 209
Appendix II-D: Pump Design ................................................................................................ 211
Appendix II-E: Compressor Design ...................................................................................... 213
Appendix II-F: Pre-Processing Design .................................................................................. 214
Appendix II-G: Amine Scrubbing Design .............................................................................. 216
Appendix II-H: Methanol Purification System Design ......................................................... 224
Appendix II-I: Quench Tank Design...................................................................................... 228
Appendix II-J: Heat Exchanger Design ................................................................................. 230
Appendix II-K: Reactor Heat Exchanger Design ................................................................... 235
Appendix II-L: Reactor Design.............................................................................................. 236
Appendix II-M: Solar Field Costing....................................................................................... 241
Appendix II-N: Cyclone Costing ........................................................................................... 242
Appendix II-O: Aspen PLUS Separator Costing .................................................................... 243
Appendix II-P: Pump Costing ............................................................................................... 244
Appendix II-Q: Compressor Costing..................................................................................... 247
Appendix II-R: Pre-Processing Costing................................................................................. 248
Appendix II-S: Amine Scrubbing Costing ............................................................................. 250
Appendix II-T: Methanol Purification System Costing ......................................................... 252
5
Appendix II-U: Quench Tank Costing ................................................................................... 253
Appendix II-V: Heat Exchanger Costing ............................................................................... 254
Appendix II-W: Reactor Heat Exchanger Costing ................................................................ 256
Appendix II-X: Reactor Costing ............................................................................................ 257
Appendix III: Computer Process Modeling and Simulations ................................................... 258
Appendix III-A: Biomass Pre-Processing Simulation ............................................................ 258
Appendix III-B: Aspen PLUS Simulation ............................................................................... 259
Appendix III-C: Amine Scrubbing Simulation ....................................................................... 260
Appendix III-D: Methanol Purification Simulation .............................................................. 261
Appendix III-E: Heat Integration .......................................................................................... 262
Appendix III-F: Reactor Computer Code .............................................................................. 263
Appendix IV: Economic Analysis Spreadsheets ....................................................................... 266
Appendix IV-A: Base Case without Carbon Credit ............................................................... 266
Appendix IV-B: Base Case with Carbon Credit..................................................................... 274
Appendix IV-C: Summary of Utilities ................................................................................... 282
6
I. Executive Summary
As carbon emissions become a growing cause for global concern, greater pressure has been
placed on industry to develop innovative alternatives to traditional commodity chemical
production. In order to investigate such an alternative, a design report has been written
examining the construction and economic feasibility of a Solar-Thermal Biomass Gasification
facility. This facility will serve as an alternative means of high-purity, industrial scale methanol
production.
The facility modeled here utilizes 204 million pounds of corn stover biomass per year as feed
stock, employs 111 full-time operators, and produces 58,300,000 gal/year of methanol end
product. The plant operates in five distinct subunits. Waste corn stover enters the biomass preprocessing portion of the facility where it is ground into usable cellulose and lignin. The usable
biomass is then sent to the biomass gasification subsystem, in which a series of three reactors
convert the biomass to methanol. In order to mitigate the environmental impact and utility
costs of the largest reactor, a solar field operating as part of the facility supplies thermal energy
to the solar reactor. An amine scrubbing system purifies the waste gas stream of environmental
toxins, while the final stage of product processing entails the purification of the end product
methanol, resulting in a final product stream with 99.97% purity by weight. The capital cost of
the facility was determined to be $300.5M.
An economic analysis was performed for plant operation in which 12.5% fixed IRR was
stipulated for facility investors. This economic analysis returned a 10.8% ROI, 9.2 year PBP and
$62.462M NPV based on a 30-year expected facility lifespan with a single year construction
period and single year of 50% capacity startup operation. In order to obtain the required 12.5%
IRR, the final product selling price was determined to be $1.69/gal methanol. This price is not
competitive with the current commodity market value of $1.05/gal (Methanex, 2015). Because
of the fa ilit s i a ilit to e su e i esto s suita le etu s hile
eeti g e d-product market
value, it is the recommendation of this design team that the Solar-Thermal Biomass Gasification
facility not be constructed. In the event that a carbon credit is granted to the facility to
incentivize eco-forward industry, a subsidy of $0.21/lb CO2 avoided would be required to
reduce the product selling price to market value and render the project economically viable.
7
II. Introduction
Project Description and Scope
Mission Statement
United States methanol consumption is on the rise and is expected to increase 26% by 2020
(The American Oil and Gas Reporter, n.d.). This opens the market for increased demand in
chemicals, transportation, and power generation, as methanol is a key commodity in all three
sectors. The chemical industry uses this versatile compound in hundreds of chemicals including
solvents, plastics, paints, and adhesives. Combustion of fossil fuels, namely petroleum based
products, in the transportation sector is the second largest source of CO2 emissions in the US
and accounts for a third of all greenhouse gas (GHG) emissions to the atmosphere (EPA, n.d.).
Alternative fuel sources that result in significantly less GHG emissions than conventional fuel
has e o e a e essit as the o ld s populatio a d e o o
o ti ues to i
ease.
Methanol provides an attractive alternative fuel option to replace petroleum due to a variety of
advantages. Implementing methanol into the transportation fuel industry could help to
significantly reach federal and state carbon reduction goals. In addition, companies are
exploring ways to use methanol as an additional fuel source for power generation to drive
turbines and create electricity.
This design project proposes to develop a renewable methanol production plant in Daggett, CA.
A techno-economic analysis will be performed to ascertain the viability of the plant in terms of
health and safety, equipment design, and return on investment. The plant will utilize waste
biomass as the feedstock and subsequently produce 58.3 million gallons of 99.97% pure
methanol annually. Slight excess of methanol was produced to account for unforeseen major
mainte a e issues that ould ut i to p odu tio ti e. I this
a
e , the u e s suppl
ould ha e li ited i te uptio s o e the ou se of the pla t s lifeti e, if a
at all. The
gasification of the biomass will be operated using a hybrid reactor with energy sources from
o e t ated sola po e a d atu al gas, hi h ill lo e the pla t s o e all GHG e issio s.
Process Description
This process is utilizing waste biomass, corn stover, in a thermochemical gasification reaction to
produce methanol. The plant will be located in Daggett, CA and will operate 24 hours a day, for
8
a total of 8000 hours per year. The plant can be divided into pre-processing, gasification, and
purification sections. Pre-processing involves drying the corn stover feedstock and grinding into
small particles applicable for gasification. Corn stover is delivered to the plant after harvest and
is initially shredded to reduce the bulk size to no larger than 6mm. The feedstock is sent to a
direct-contact air-dryer to convectively pull moisture from the particles. A hammer mill is then
employed to reduce the particles to micron size. Lastly, a hopper is used to pressurize the feed
to 35 bar to meet the specifications of the solar reactor.
At this point the feedstock is at the correct temperature and pressure for the gasification
reaction. Because biomass reactions are complicated and novel, software such as Aspen PLUS
requires multiple pieces of equipment to model them, though the reaction would take place in
a single vessel. The gasification reactor has concentrated solar power and natural gas as the
energy sources. Concentrated solar power (CSP) is produced by a field of heliostats targeted at
a solar tower with a compound parabolic concentrator. CSP is one of few renewable
technologies able to achieve the high temperatures required in the gasification process
(>1000°C) due to concentrating the thermal energy. Natural gas is fed as a supplement to CSP
to allow 24 hour operation.
In the gasification step, the pre-treated biomass must be reacted in high temperature, lowoxygen conditions with water and a methane stream. A controlled level of oxygen limits the
combustion reaction formation of carbon dioxide, but provides enough heat for subsequent
gasification reactions. High temperatures yield fewer hydrocarbons such as char, tar, and ash
and increases conversion directly to carbon monoxide and hydrogen (syngas). The addition of
water and methane also serves to reach to the desired 2:1 ratio of H2:CO in the syngas and to
diminish the selectivity of carbon dioxide formation. The chief gasification reactions are shown
in the equations below:
�
→[
+
+
+
+
+
→
]+
→
+
+
+
ℎ
(1)
(2)
(3)
9
+
+
+
→
+
+
→
→
→
→
(4)
(5)
+
(6)
+
(7)
(8)
The gasification process produces dirty, contaminated gas streams that undergo various
separation processes to clean the gas and remove particulates, such as ash solids, acid-gas,
sulfur, and chlorine. The resulting vapor stream is sent to the methanol processing reactor.
Maximum production of methanol follows from a high concentration of CO, a low
concentration of CO2, and an optimal ratio of H2:CO:H2O in the feed stream, values that are
dependent on specific reactor conditions. Equations (9), (10) and (11) below summarize the
simultaneous methanol synthesis reactions:
+
+
+
→
→
→
+
(9)
(10)
+
(11)
The output stream from the methanol reactor is split into a recycle stream back to the solar
reactor, a purge stream that is flared, and a product stream sent downstream for purification.
The downstream methanol purification process operates 24 hours a day as well. The dirty
methanol stream employs distillation to obtain a final methanol product stream with 99.97%
purity.
10
III. Background Information
Market for Renewable Methanol
The $36 billion methanol industry has 90+ plants in operation worldwide and produces 49.8
million tons annually to be used in the chemical, transportation and power generation sectors
(Methanol Institute, n.d.). Methanol is the simplest alcohol, with a chemical formula of CH3OH.
Its simplicity brings forth a variety of characteristics; methanol occurs naturally in the
environment, it is biodegradable, light and colorless, and it quickly breaks down in aerobic and
anaerobic conditions (Methanol Institute, n.d.). Methanol can be produced from a diverse array
of feedstocks which gives it the distinct advantage of polygeneration: the ability to be made
from any resource that can be converted into synthesis gas. As a result, methanol is classified as
o e tio al o
e e a le. Conventional methanol is produced from steam reformation
with the use of a fossil fuel, coal or natural gas, and steam. This produces a synthesis gas, as
shown below in Equation (12) for steam reformation (Methanol Institute, n.d.).
+
→
+
+
→
+
ℎ �
(12)
+
(13)
The syngas is then processed over a catalyst to yield CH3OH, as shown in Equation (13).
+
+
Renewable methanol is produced from the synthesis of waste biomass such as switchgrass,
forest trimmings, corn stover, or other agricultural residue products in a very similar manner of
gasification and catalysis. Using biomass as the feedstock is considered carbon neutral due to
the sizeable carbon absorbance of the feedstock before processing. Additionally, renewable
biomass has an advantage in that it looks to include a CO2 stream as the feedstock,
consequently utilizing a GHG in its process (Karen Law, 2013).
Methanol is an exceptionally diverse chemical and thus has many advantages as a material for
chemical production or as a fuel source. With its unique ability for polygeneration,
manufacturers can tap into multiple resources to supply the increasing methanol demand.
Since methanol production is already well established in the global marketplace, the existing
infrastructure and logistics would allow for production conversion between conventional and
11
renewable means, and possibly allow for a full transition to renewable methanol in the future.
As well, the US and other countries can leverage the economic benefit of the increased demand
of methanol. This is seen in two recent development projects: a 10 year Chesapeake Energy
Corp-Methanex contract and G2X Energy methanol-to-gasoline plant construction (The
American Oil and Gas Reporter, n.d.). Collaborations with natural gas energy companies and
new plant development projects represent a growing market that is seeking ways to utilize the
full potential.
At 40% of supply, the chemical sector has the largest demand for methanol, which is usually
provided through conventional means. The largest chemical use fo
etha ol is i the plasti s
industry to produce resins and polymers (The University of York, n.d.). Common methanol
derivatives include dimethyl ether, formaldehyde, acetic acid, methyl methacrylate, and
methylamines.
In the
sa d
s,
etha ol e ei ed atte tio as a alte ati e fuel sou e, a el to
create a fuel blend with gasoline and possibly ethanol. It did not become a substantial
commodity though, primarily due to falling petroleum prices which negated any economic
incentive to its usage. In recent years, as petroleum prices and supply has been more volatile,
many countries including the United States are seeking ways to gain energy security and
mitigate uncertainties. As a result, the US has a seen a steady rise in alternative fuel demand
over the last decade. Federal Renewable Fuel Standards (RFS) calls for 18.11 billion gallons of
alternative fuel production in 2016; a 7% increase over the 2015 standards (EPA, n.d.).
Additionally, the Open Fuel Standard bill, if approved, would greatly promote the need for
methanol and demand could increase exponentially (Open Fuel Standard, n.d.) Methanol is
seen as one of the chief commodities able to fill the demand gap in alternative fuels due to a
variety of advantages specifically in the transportation sector: studies have shown a 65-95%
carbon reduction from well-to-tank and a 15-20% lower tailpipe emissions with fuel blends of
methanol, it cuts nitrogen oxide and volatile organic compound emissions, there are no toxic
additives necessary and has a half-life of 1-7 days versus gasoline and other constituents at over
a hundred days (Methanol Institute, n.d.). Another significant advantage is its ability to be used
in advanced efficiency technologies such as PEM fuel cells. While there are considerable
12
advantages, methanol use in transportation fuels has disadvantages as well, such as it has half
the energy content as gasoline on a volume bases and it is corrosive in nature and miscible with
water which places different material requirements in vehicles (Methanol Institute, n.d.).
Lastly, methanol is being explored as an emerging fuel for electrical power generation. It can be
used as an energy carrier for hydrogen storage and delivery. Research has shown it can be
incorporated into existing dual-fueled gas turbines and can stand alone as a turbine fuel
(Methanol Institute, n.d.).
Renewable Methanol from Biomass
As opposed to bioethanol, methanol production from biomass has very few limitations in
feedstock choice. Agricultural resides, wastes, forest trimmings, and wood can be used – any
material able to undergo gasification to produce syngas. Since the feedstock can be a waste
material, there is little competition between materials usable as food sources. Furthermore, the
Energy Independence and Security Act mandates that non-food based bio-fuels ramp
production to meet the goal of 36 billion gallons by 2022, where methanol can obviously play a
major role.
A disadvantage to using biomass is that several additional production steps need to be taken
into consideration. The biomass must be pretreated by conditioning and drying to break down
cellulose and lignin and densify the material. Additionally, the use of biomass has a lower H:C
ratio compared to natural gas and thus produces char and ash to be taken into consideration in
the separations and waste disposal process. Also, conditions in the syngas formation reactions
eed to e
a ipulated fo
a i u
o e sio . The s gas eeds to e upg aded to
increase the H2 content and lower the methane content by utilizing the water-gas shift
reaction. Overall, biomass lowers the efficiency of the overall process due to these factors,
though methanol selectivity over 99% can still be achieved (Cheng, 2010).
Many feedstock options are available, though corn stover is the most plentiful non-food, noncrop source of biomass in the US, due to corn crops producing the highest volume of residue in
comparison to all other major crops (Dupont, 2012). The current supply has been estimated at
75 million tons per year (Roth, 2014). As well, due to the increase in demand of corn and higher
13
crop yields, the level of corn residue has increased as well, though the removal must be
sustainable.
Solar-Thermal Processing
Solar energy is one of the most accessible forms of renewable energy, and the amount available
is more than enough to power the entire world, provided it can be harnessed (Clean Technica,
n.d.). Currently, the capacity in the United States is 1.75 operating GW, with 2.2 GW in
development. Globally, there are 4.7 GW total current operating capacity with 22GW expected
by 2025 (Heba Hashem, 2015) . Among the solar energy possibilities, concentrated solar power
(CSP) has emerged as the major conversion technology due to many of its unique features and
relatively high efficiencies. CSP has the capability to achieve extremely high temperatures and
can be integrated with other conventional fossil fuel plants to create hybrid systems.
Consequently, a traditional fossil fuel pla t s o e all GHG e issio s a e lo e ed
edu i g
the fossil fuel input and relying partially on solar. CSP also has the ability for thermal energy
storage which allows for extended operation beyond when the sun is radiating energy.
A CSP process is constructed to collect the sunlight using heliostat mirrors that concentrate
solar energy using compound parabolic concentrators (CPC) to a centrally located tower. This
type of concentration allows for temperatures greater than 1000°C, which is beneficial to many
reaction systems. Figure 1 shows an aerial view of a solar power tower and heliostat field.
Figure 1: Aerial view of solar power tower and heliostat field
14
In the biomass to methanol plant, a renewable energy source was sought to mitigate the use of
natural gas and GHG emissions. As well, biomass gasification requires high temperatures in
order to avoid tar formation and to increase the conversion to higher energy gases such as CO
and H2. CSP was the obvious choice to meet both the high temperature demands and the
lowered GHG emissions.
15
IV. Environmental, Health & Safety
As in any chemical manufacturing system, it was paramount to consider the effects of the
proposed biomass gasification and methanol production plant on the environment and the
health and safety of the operators. This process involved a variety of chemical components
including reagents, catalysts, and waste products that are toxic, flammable, or highly reactive.
Waste treatment or management was considered for all streams exiting from the process. In
addition, the simulated process operated at temperatures over 1000°C and pressures up to 80
bar, conditions which can cause serious harm or disaster if not well-controlled and monitored.
Accordingly, worst-case scenarios were developed and evaluated. Finally, a life cycle analysis
(LCA) was performed to evaluate the environmental impacts of the proposed process across its
30-year lifetime on public health and land.
Chemical Hazards
A summary of all the chemical components present in the system is presented in Table 1 below.
Included in this table are the chemical formulas, lower explosion levels (LELs) of each
component, their auto-ignition temperatures, important safety hazards, general safety hazards,
and the permissible exposure limits (PELs) established by the Occupational Safety and Health
Administration (OSHA) (OSHA, 2011). In the following sections, methodologies to reduce these
chemical hazards will be discussed.
Table 1: Overview of chemical components in the biomass gasification and methanol production process. OSHA PELs are
obtained from (OSHA, 2011) or the indicated source (material safety data sheet, MSDS). GHG = greenhouse gas. *OSHA PELs not
considered because the limiting safety factor is the available oxygen in the atmosphere.
Component
(phase)
Chemical
formula
Biomass fly ash
(s)
Hydrogen (g)
N/A
Lower
explosion
level
N/A
H2
4 vol%
Auto-ignition
temperature
(°C)
Wide range
(232-2760)
565.5
Nitrogen (g)
N2
Not
flammable
Not
flammable
Water (g, l)
H2O
Not
flammable
Not
flammable
Safety and
environmental
hazards
Irritant; solid
hazardous waste
Flammable,
asphyxiant*; none
Asphyxiant*;
none
Slipping; none
OSHA
PELs
(ppm)
6
-
None
Source
(Weyerhaeuser,
2014)
(Air Products,
1994)
(Air Products
and Chemicals,
Inc., 1997)
(Sciencelab.com,
Inc., 2013)
16
Corn stover
biomass (s)
Multicomponent
25 g/m3
Hydrogen
sulfide (g)
H2S
4.3 vol%
Dust layer:
215
Dust cloud:
450
270
Methane (g)
CH4
1.8 vol%
287
Carbon
monoxide (g)
CO
≥
700
Hydrochloric
acid (g, aq)
HCl
Not
flammable
Not
flammable
Carbon dioxide
(g)
CO2
Not
flammable
Not
flammable
Ethane (g)
C2H6
2.9 vol%
N/A
Ethylene (g)
C2H4
2.7 vol%
450
Nitric oxide (g)
NO
Not
flammable
Not
flammable
Nitrogen
dioxide (g)
Ammonia (g,
aq)
NO2
Not
flammable
16 vol%
Not
flammable
651
Zinc oxide (s)
ZnO
Not
flammable
Not
flammable
Zinc sulphide
(s)
ZnS
Not
flammable
Not
flammable
Zinc chloride
(s)
ZnCl
Not
flammable
Not
flammable
Methyl
diethanolamin
e (MDEA) (l,
aq)
Mineral oil
CH3N(C2H4O
H)2
1.4 vol%
280
Varies
N/A
N/A
NH3
.
ol%
Explosive, irritant;
none
6
(IEA Bioenergy,
2013)
Flammable, toxic;
toxic to aquatic
life
Flammable,
asphyxiant*; GHG
Flammable, toxic;
lethal to fish,
pollutant
Corrosive, toxic,
irritant; acidifies
water
Asphyxiant; GHG
10
(Airgas, 2015)
-
(Airgas, 2015)
Flammable,
asphyxiant*; none
Flammable,
asphyxiant*; none
Toxic, oxidizer,
irritant, reactive;
pollutant
Toxic, carcinogen;
pollutant
Flammable,
corrosive, irritant;
toxic to aquatic
life, pollutant
Toxic, irritant;
hazardous solid
waste
-
Toxic, irritant,
reactive;
hazardous solid
waste
Toxic, corrosive,
irritant;
hazardous solid
waste
Irritant;
biodegradable,
low toxicity to
aquatic life
Irritant, slightly
flammable;
hazardous waste
25
(Matheson TriGas, Inc., 2008)
0.3
(Sciencelab.com,
Inc., 2013)
5000
(Air Products
and Chemicals,
Inc., 1994)
(Sigma-Aldritch,
2015)
(Airgas, 2015)
25
(Matheson TriGas, Inc., 2008)
1
(Matheson TriGas, Inc., 2008)
(Air Products
and Chemicals,
Inc., 1999)
25
5 (fume),
6
(particulat
e)
N/A
(Sciencelab.com,
Inc., 2013)
1 (fume),
6
(particulat
e)
N/A
(Sciencelab.com,
Inc., 2013)
5 (mist)
(Sciencelab.com,
Inc., 2013)
(Sciencelab.com,
Inc., 2013)
(Union Carbide
Corporation,
2015)
17
From this table and the accompanying MSDS documents, the most serious chemical hazards
were evaluated. Many of the chemical components in this process were present in the gas
phase, posing hazards regarding the availability of oxygen in the atmosphere on-site. In
particular, methane, ethane, ethylene, nitrogen, hydrogen, and carbon dioxide are classified by
OSHA as simple asphyxiants. A simple asyphxiant poses a hazard when the amount of available
oxygen in the air drops below 10%, which can cause unconsciousness and, in the absence of
further action, death by suffocation. If a large process gas leak occurred, this chemical hazard
would be a concern.
In addition, many of the components involved in the gasification of biomass are highly
flammable: hydrogen, carbon monoxide, hydrogen sulfide, methane, ethylene, ammonia, and
the biomass itself. Since the process gases are at pressures of at least 35 bar, it is unlikely that
ambient air would leak into the process. If the gas were exposed to ambient air at the
simulated temperatures up to 1450°C, the gas would likely auto-ignite, causing flaming jets of
high temperature, high pressure gas to be ejected from the equipment. These jets could start
fires in other equipment elsewhere in the plant, severely injure operators, and potentially cause
an explosion in other high pressure equipment. In addition, there is a risk of biomass fires
occurring. After the corn stover used in the process is unloaded, it is likely to be stored in a silo
before pre-processing. Since the biomass is not already dried (25% moisture by mass), there
exists a possibility of self-heating by microbial heat generation and exothermic side reactions in
ambient air. High temperatures can result at the core of a biomass silo as a result of thermal
runaway, ultimately resulting in spontaneous combustion of the biomass material (IEA
Bioenergy, 2013). The off-gassing of the moist biomass can also produce flammable compounds
such as CO, CH4, and aldehydes. The storage silo can then explode as a result of this volatile
mixture reaching high temperatures; an example of a silo explosion can be seen in Figure 2
below.
18
Figure 2: Silo explosion caused by the ignition of biomass off-gasses (IEA Bioenergy, 2013)
Many of the components in this process are also harmful to human health, particularly
ammonia, carbon monoxide, nitrogen oxide, nitrogen dioxide, and hydrogen sulfide. When
ammonia is inhaled, it is severely irritating to the lungs. The chemical is also corrosive to the
skin and eyes as a gas and an aqueous solution, where it contributes to an overall alkaline
solution. Upon inhalation, carbon monoxide can cause suffocation, blood damage, reproductive
complications, nerve damage, brain damage, and death. Inhalation of nitrogen oxide is fatal at
low concentrations, and eye contact with the chemical causes serious eye damage. Nitrogen
dioxide burns mucous membranes in the eyes, nose, and lungs and can be fatal if too much is
inhaled. Finally, hydrogen sulfide is very toxic to pulmonary tissue and is fatal if inhaled. The
biomass can also be harmful, as workers can become exposed to bacteria, spores, or other
pathogens living in the pre-dried biomass.
O“HA sets sta da ds ega di g o ke s e posu e to ha
ful he i als i the o kpla e. Table
1 outlines the permissible exposure limits (PELs) to various chemicals, expressed in parts of
compound per million parts of air (ppm). It is also required by OSHA that all chemical containers
19
a e p ope l la elled ith the ide tit of the haza dous chemical and appropriate hazard
a i gs a d that M“D“s a e a aila le ith detailed i fo
thei effe ts, ho to p e e t e posu e, a d e e ge
atio a out he i al haza ds,
t eat e t if a e posu e o u s
(OSHA, 2014). Finally, an inventory of chemicals present at the facility must be kept current and
available for reference. Personal protective equipment (PPE) is required for workers exposed to
these chemicals – these specifics will be addressed in the following section.
Health and Safety Considerations
Operator Safety
In order to legally employ operators, engineers, and other staff to run the plant, OSHA dictates
that certain minimum safety standards in plant operation and design be met. These regulations
are in place to protect employees from hazards in the workplace, as well as provide protections
to o ke s i
epo ti g these haza ds. It is the e plo e s responsibility to provide certain
information to meet OSHA standards (OSHA, 2014). Injury and illness records must be available
to employees if these injuries or illnesses are a result of workplace conditions. Workers have
the right to exposure data; that is, employers must monitor levels of chemicals or substances
regulated by OSHA PELs and provide this data to employees. Finally, workers have the right to
thei
edi al e o ds, espe iall if a o ke s health has ee affe ted e ause of e posu es
at o k (OSHA, 2014).
In case of an evacuation of the plant because of a disaster or other safety concern, exit routes
must be provided and meet certain OSHA requirements according to the Code of Federal
Regulations Title 29 (CFR 29), standard 1910.36 (OSHA, 2012). Exit routes must be permanent
and be built out of fire- esista t
ate ials. The e it
ust e protected by a self-closing fire
door that remains closed or automatically closes in an emergency upon the sounding of a fire
alarm or employee alarm system (OSHA, 2012). The exit must also be unlocked and have
direct, unobstructed access to an open space. For example, the open space can be a street,
alley, or walkway. Outdoor exit routes are permitted under standard 1910.36.
Occupational noise exposure is an important workplace condition that must be met in industrial
settings; CFR 29, standard 1910. 95 outlines some requirements (OSHA, 2012). Table 2 presents
20
permissible noise exposures in duration per day versus sound level in decibels (dB). Proper
engineering and PPE should be provided if workers are exposed to sound exceeding these
values and durations. Examples of proper PPE include earplugs and noise-cancelling earmuffs.
Noise levels must also be monitored with calibrated audiometers and the data recorded.
Hearing protectors must also be provided free of charge to the workers.
Table 2: Permissible noise exposures (OSHA, 2012)
Duration per day, hours
Sound level, dB
8
90
6
92
4
95
3
97
2
100
1 1/2
102
1
105
1/2
110
1/4 or less
115
OSHA also regulates the quality and condition of working-walking surfaces within industrial
plants, according to CFR 29, standard 1910.22 (OSHA, 2012). These surfaces must be kept clean
and dry in order to prevent slips and falls by workers. If the area surrounding process
equipment is wet, drainage systems must be installed and dry standing areas must be
maintained nearby. Aisles and passageways must be kept clear of debris and clutter. On
platforms and ladders, guard rails must be installed to prevent employees and operators from
falling from great heights.
For equipment operating at high pressures, OSHA requires the installation of pressure relief
valves according to CFR 29, standard 1910.101 (OSHA, 2012). These safety measures are
implemented to prevent the explosion of vessels in the case of overpressurization.
Overpressurization may result from blockage in a line or a larger temperature than expected in
a process unit. Pressure monitors equipped with alarms should also be installed in the process
to alert operators to rapidly rising pressures. Since this process also uses a variety of flammable
21
chemicals with a possibility for fires, a fire suppression system should be installed throughout
the plant and within the biomass silos to prevent flames.
Finally, personal protective equipment (PPE) must be worn by operators at all times to prevent
harm by chemicals, equipment, or operating conditions in the plant. Biomass or fly ash dust
clouds may form in the plant, so respirators should be worn at all times. Respirators also
prevent the inhalation of solid hazardous waste such as spent zinc oxide catalyst. If there is a
process gas leak, self-contained breathing apparatuses should be used to prevent the inhalation
of potentially fatal chemicals. Gloves and safety glasses should be worn at all times to prevent
injuries to the hands and eyes and block these bodily areas from discharges of chemicals.
Chemical goggles should be used upon if there is a process gas leak or if handling pure
chemicals or hazardous waste. Protective clothing such as overalls and long sleeves should be
worn at all times to prevent chemical exposure and skin injuries. A hard hat should be worn at
all times to prevent head injuries from falling equipment. Finally, to prevent the formation of
sparks from static charge, anti-static clothing and conductive boots should be worn at all times.
With these plant design and PPE enforcements in place, operator safety is enhanced.
Licensure and Permits
In order to operate a chemical processing facility in unincorporated San Bernardino County, CA,
a collection of licensures, permits, and information must be obtained from a variety of federal,
state, and county sources. These specifics are outlined in Table 3 below. A Certificate of
Disclosure of Hazardous Substances must be obtained from the Certified Unified Program
Agency (CUPA), Hazardous Materials Division within the San Bernardino County Fire
Department since hazardous materials are on site; this document is also known as a Business
Emergency/Contingency Plan. CUPA also provides hazardous materials and waste generation
information. Fire prevention information is supplied by the County of San Bernardino; this
authority will provide an inspection of the plant to ensure fire suppression and control systems
are being maintained effectively. The Air Quality Management District (AQMD) provides
Authority to Construct, Permit to Operate, and Building permits to industrial projects that emit
air emissions in the county. AQMD also provides Certificates of Occupancy. Since preliminary
hazardous waste treatment is done in the proposed plant, a Hazardous Waste Facility Permit
22
will be required from the Department of Toxic Substance Control. A State EPA ID Number is
e ui ed f o
the Depa t e t of To i “u sta es Co t ol fo
usi esses that generate,
surrender to be transported, transport, treat, or dispose of hazardous waste (California
Governor's Office of Business and Economic Development [GO-Biz], 2015). Since wastewater
will be discharged after treatment, an Industrial Activities Storm Water General Permit must be
obtained from the California Environmental Protection Agency (Cal/EPA). Cal/EPA also sets
Waste Discharge Requirements (WDRs). The Department of Industrial Relations provides
Occupational Safety and Health Information for businesses to use to develop an Injury and
Illness Prevention Plan. The Depa t e t of I dust ial ‘elatio s also p o ides Wo ke s
Compensation Information. The Employment Development Department provides registration
forms for employers that pay over $100.00 in wages to one or more employees. The Franchise
Tax Board provides state income tax information and forms. An Employer Identification
Number (EIN) or Social Security Number (SSN) is required by the Internal Revenue Service (IRS)
in the U.S. Department of Treasury for all employers for income tax purposes. Employees are
required to submit Proof of Residency forms to demonstrate proof of eligibility to work in the
U.S. Finally, a Title V Air Permit is required for any facility that emits large quantities of nitrogen
oxides, or operates in a state subject to federal Acid Rain regulations. (Environmental
Protection Agency [EPA], 2015).
Table 3: Summary table of permits, licenses, and information eeded fo a Che i al o Pai t Fo ulatio
usi ess i
unincorporated San Bernardino county, California (California Governor's Office of Business and Economic Development [GO-Biz],
2015)
Licensure, permit, or information needed
Certificate of Disclosure of Hazardous
Substances
Hazardous Materials/Waste Generation
Fire Prevention Information/Inspection
Authority to Construct/Permit to Operate
Certificate of Occupancy/Building Permit
Hazardous Waste Facility Permit
State EPA ID Number
Industrial Activities Storm Water General
Permit
Waste Discharge Requirements
Distributing authority
San Bernardino County Fire Department
Level of government
County
San Bernardino County Fire Department
County of San Bernardino
Air Quality Management District
Air Quality Management District
Department of Toxic Substances Control
Department of Toxic Substances Control
California Environmental Protection
Agency
California Environmental Protection
Agency
County
County
Regional
Regional
State
State
State
State
23
Occupational Safety and Health
Information
Workers' Compensation Information
Registration Forms for Employers
State Income Tax Information
Employer Identification Number (EIN or
SSN)
Proof of Residency Requirement
Title V Permit
Department of Industrial Relations
State
Department of Industrial Relations
Employment Development Department
Franchise Tax Board
U.S. Department of Treasury
State
State
State
Federal
U.S. Immigration and Naturalization
Service
California Environmental Protection
Agency
Federal
Federal?
Environmental Considerations
One of the largest goals in the proposed biomass gasification for methanol production process
was to utilize a renewable energy source to provide the high temperatures needed in the
gasification reactor rather than burn fossil fuels. However, there were other environmental
considerations that were taken into account to evaluate the environmental impact of the plant,
including individual chemical considerations, possible chemical spills, greenhouse gases, and
waste disposal. Greenhouse gas emissions over the lifetime of the plant were considered in a
life cycle analysis (LCA).
Some of the chemicals given in Table 1 are listed as hazardous by the California Environmental
Protection Agency (Cal/EPA). These hazardous materials were classified as characteristic
hazardous wastes or used oils. The characteristic wastes are classified as such if they exhibit
a
of the fou
to i it
ha a te isti s of a haza dous aste ig ita ilit , o osi it , ea ti it , a d
a d a e gi e i Table 4 (California Department of Toxic Substances Control, 2014). An
ignitable substance is one that can cause a fire, spontaneously combust, or has a flash point less
than 60°C. Corrosive substances are materials that produce acidic or alkaline solutions (pH < 2
o pH >
.
o
o ditio . To i
o ode
etal sto age o tai e s. ‘ea ti e astes a e u sta le u de
ate ials a e ha
e tai
ful o fatal [to fish o hu a s] when ingested or
a so ed. Fi all , used oil efe s to a
oil that has ee
efi ed f o
ude oil, o a
synthetic oil that has been used and, as a result of use, is contaminated with physical or
he i al i pu ities (California Department of Toxic Substances Control, 2014).
24
Table 4: Classification of hazardous waste components in the biomass gasification for methanol production process. Chemicals
from Table 1 are classified according to their characteristic waste codes or their status as a used oil.
Chemical
(I)
(R)
(C)
(T)
Used
oil
Biomass fly ash
Hydrogen
Nitrogen
Water
Corn stover biomass
Hydrogen sulfide
Methane
Carbon monoxide
Hydrochloric acid
Carbon dioxide
Ethane
Ethylene
Nitric oxide
Nitrogen dioxide
Ammonia
Zinc oxide
Zinc sulphide
Zinc chloride
MDEA
Mineral oil
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
X
After classifying these streams as hazardous waste, the waste disposal of these streams was
considered. A table of all of the outlet streams from the process is given in Table 5 below. Two
of these streams were immediately flared: Waste Vap and PURGE. These streams are purged
because the products of their combustions are not hazardous. Two more streams require some
treatment before being flared: Light HC and Acid Gas. These streams will be passed over zinc
oxide (ZnO) to react all the hydrochloric acid (HCl) and hydrogen sulfide (H2S) according to
Equation (79) as presented in the Zinc-Oxide Reactor equipment design section, then the
streams are flared. In this way, harmful sulfur oxide (SOx) emissions can be avoided, and HCl will
not be emitted to the atmosphere. The amount of ammonia to be flared is small compared to
the outlet mass flow rate to be flared, and syngas burns with a lean flame and a combustion
speed
u h ui ke tha
atu al gas (U.S. Department of Energy, 2006). So, it was
determined to be acceptable to flare ammonia, despite the nitrogen oxide emissions.
25
Table 5: Summary of all outlet streams from the proposed process with compounds that are hazardous, sorted by simulation.
Components with mole fractions less than 1E-7 are not included.
Stream
Simulation
Mass flow rate
(lb/hr)
Composition (wt%)
Phase
Waste disposal
method
Bottoms
HYSYS (MeOH)
6037
H2O: 0.9524
MeOH: 0.0476
Liquid
Modified LuckEttinger process,
then discharged as
wastewater
Waste Vap
HYSYS (MeOH)
126.5
H2: 0.0193
CH4: 0.0363
C2H4: 0.0001
N2: 0.0483
H2O: 0.0125
CO: 0.3085
CO2: 0.0024
MeOH: 0.5725
Vapor
Flared
Light HC
HYSYS (Acid Gas)
1.316
H2O: 0.0117
CO2: 0.1359
H2S: 0.0015
CH4: 0.0087
CO: 0.7370
H2: 0.0984
N2: 0.0068
Vapor
Passed over ZnO,
then flared
Acid Gas
HYSYS (Acid Gas)
1317
H2O: 0.6201
CO2: 0.3729
H2S: 0.0046
CO: 0.0011
NH3: 0.0012
H2: 0.0001
Vapor
Passed over ZnO,
then flared
Purge
HYSYS (Acid Gas)
616.9
MDEA: 0.4500
H2O: 0.5262
HCl: 0.0238
Liquid
Passed over ZnO,
then recycled into
amine scrubbing
system
QNCH-H2O
Aspen PLUS
59227.1
All H2O
Liquid
Released as steam to
the atmosphere
SLDWASTE
Aspen PLUS
226.878
All Ash
Solid
Sold
AQ-WASTE
Aspen PLUS
2141.34
H2O: 0.99915
H2S: 1.3E-6
NH3: 0.00084
CO2: 4E-6
Liquid
Modified LuckEttinger process,
then discharged
ACIDS
Aspen PLUS
See HYSYS
26
ZN-SOLID
Aspen PLUS
0.336905
ZnS: 1.0000, trace
ZnCl
Solid
Solid hazardous
waste
PURGE
Aspen PLUS
3141.47
N2: 0.05385
MeOH: 0.04060
C2H4: 0.00002
H2O: 0.00145
H2: 0.11288
CO: 0.77194
CO2: 0.00029
CH4: 0.01896
Vapor
Flared
The final waste products zinc chloride (ZnCl) and zinc sulfide (ZnS) are disposed of as solid
hazardous waste. The solid waste stream ZN-SOLID is also disposed of as solid hazardous waste.
The amount of ZnO required to react with the HCl and H2S in each of solid waste streams was
determined by assuming 100% conversation of the gases and is given in Table 6. This analysis
yielded a total amount of 112,612 kg/yr needed of ZnO and 163,206 kg/yr of solid waste for
disposal. This solid waste will be placed into a satellite accumulation area on-site and routinely
taken to a waste disposal and treatment center (EPA, 2012). The waste solid ash from
SLDWASTE can be sold as fly ash to a concrete producer for $0.012/lb (The Aberdeen Group,
1985). The excess quench water from the QNCH-H2O stream will be emitted as steam to the
atmosphere.
Table 6: ZnO catalyst needed for waste treatment. ZnO costs $0.20/kg, and it costs $0.31/kg of solid waste for disposal. Assumes
8000 hr/yr of operation.
Stream
Mass flow rate of ZnO needed
Mass flow rate of ZnS +
Cost of disposal of ZnO
(kg/hr)
ZnCl2 (kg/hr)
and waste disposal ($/hr)
Light HC
0.004
0.004
0.002
Acid Gas
6.51
7.80
3.72
Purge
7.43
12.45
5.35
ZN-SOLID
0.128
0.153
0.073
TOTAL
14.1
20.4
9.14
112,612
163,206
73,120
TOTAL (per yr)
To treat the remaining wastewater, a single-process wastewater treatment process called the
Modified Luck-Ettinger (MLE) process will be used (Exponent, 2012). MLE is a simple process
that utilizes two steps to remove nitrogen from wastewater – a simplified schematic is given in
27
Figure 3 below. The first step is the feeding of nitrogenated wastewater (AQ-WASTE) to an
alkaline anoxic tank, where bacteria take ammonia and oxidize it to nitrates; this process is
called denitrification. The bacteria needed reside in municipal waste, which is assumed to be
readily available from Daggett nearby. In MLE, an additional carbon source is needed; in this
process, the methanol waste stream Bottoms supplies this additional carbon. This decision
makes sense because the methanol is relatively dilute, reducing the toxicity of the methanol to
the bacteria. These nitrates and the added methanol are then consumed by heterotrophic
bacteria in nitrification, which emit the reduced nitrogen from the nitrates as nitrogen (N2) gas,
which is inert and not hazardous. A clarifier then removes the waste activated sledge (WAS)
from the effluent wastewater, which is disposed as municipal waste. Some of the WAS is
recycled to the process as recycled activated sledge (RAS). The large benefit of MLE in the scope
of the proposed process is that MLE consumes some of the methanol and ammonia hazardous
waste at the same time. The overall nitrification reaction is given in Equations (14) and (15)
below. The overall denitrification reaction is given in Equation (16) below.
+
−
+
+
�
→
−
+
−
→
→
+
+
+
−
+
+
(14)
(15)
+
−
(16)
Figure 3: Simplified schematic of the Modified Luck-Ettinger (MLE) process for the denitrogenation of wastewater (Exponent,
2012).
28
Worst-Case Scenarios
When designing any chemical processing plant, it is critical to evaluate the worst-case scenarios
that could occur in order to design ways to prevent disaster. In this section, a few possible
scenarios were considered. First, a leak of high-pressure process gases could occur, as discussed
earlier. Gas monitors and alarms should be installed throughout the plant in order to detect the
presence of H2 or CO gas in the atmosphere, as this would indicate a leak in the process
equipment. From there, the plant would be evacuated to ensure safety of the operators and
engineers. If the leak occurred in a unique piece of process equipment, the process would be
shut down by defocusing the heliostat mirrors off of the solar tower and shutting off the
regenerative natural gas burner in the solar reactor. To avoid shutdown of the whole process,
redundant lines could be installed throughout the process. In the event of a leak, valves at
either end of a redundant line could be opened to divert flow away from the leaky pipe. The
leaky pipe would then be shut off from the system by closing the valves and repaired.
The quench tank serves an integral role in quenching the gasification reaction by drastically
lowering the temperature of the solar reactor effluent from 1450°C to 120°C. If cooling water
flow to the quench tank were to suddenly stop, the very hot process gas would flow further
into the process, damaging many pieces of equipment not designed to withstand temperatures
over 1000°C. To avoid this occurrence, multiple pumps of the same specifications of the one
feeding cooling water to the quench tank could be purchased and installed in parallel. One of
these pumps could be used if another were to malfunction or break. Temperature sensors
should be installed in the piping downstream from the quench tank to detect its possible
malfunction.
As mentioned before, biomass fires could occur in storage silos. Self-heating of biomass can
cause the internal temperature of the biomass to reach upwards of 400°C, causing a fire to
ignite. Figure 4 shows the progress of a silo fire over 30 hours. At 30 hours, the fire begins to be
extinguished with inert gas. This figure demonstrates the importance of placing thermocouples
at the center of the biomass silo in order to detect the higher temperatures present in a silo
fire. A fire suppression system should be installed at the bottom of any storage silo to
extinguish the biomass with inert gas in case of a fire.
29
Figure 4: Visualizatio of the easu ed te pe atu es i side a o k-up silo, 1 m diameter and 6 m height. The smoldering fire
was triggered in the middle of the silo and then allowed to develop freely which resulted in a slow fire spread downwards in the
silo. The combustion gases did not reach the top of the silo until about 20 hours. After about 30 hours, inert gas was applied
th ough the otto of the silo whi h esulted i a fast edu tio of the s olde i g i te sity (IEA Bioenergy, 2013).
Finally, the exothermic reaction present in one of the three methanol synthesis reactors could
go out of control if the cooling water to the cooling coils within the vessel were to abate. The
temperature in the system would rise, eventually causing the system to overheat and perhaps
fail. To explore this idea further, an approximate adiabatic temperature rise was calculated for
one reactor. The methanol reactors are filled 25/26 of the way with a slurry of mineral oil
(Witco-40 oil) and catalyst. If the catalyst volume is assumed to be negligible, the volume of the
reactor filled with oil
�
is calculated as follows with Equation (17):
�
where
=(
)
(17)
is the volume of all three reactors combined. The
found with using the heat capacity
and density
, �
�
�
, �
of the oil can then be
of mineral oil (Engineering Toolbox,
n.d.) (Sciencelab.com, Inc., 2013). This calculation is shown in Equation (18) below:
�
It was assumed that all of the heat duty
, �
�
�
, �
(18)
of the reactor that would normally be removed by
the cooling coils goes into the mineral oil. The
be negligible. The reactant gas
=
value of the solid catalyst was assumed to
was calculated using values from Aspen PLUS and dividing
30
the
ass flo
ate
the ea to eside e ti e τ, which was found using Excel and Polymath
during the reactor design (see Appendix III-F). This value was calculated along with
Table 7 below and was found to be 69. Therefore, the
greater than that of the gas, so the
�
, �
in
value for the oil is substantially
value of the gas can be neglected. Therefore, the
energy balance for this simplified system is given in Equation (19) as follows:
where
=
�
(19)
, �
is the temperature of the reactor and is the time after the cooling coils stop
functioning. Separating and integrating Equation (19) yields Equation (20):
where
=
=
=
+
�
(20)
, �
ℎ . A plot of Equation (20) with Excel can be seen in ___. From this plot, it
was determined that the mineral oil does a good job of diluting any temperature runaway that
may occur in the operation of this reactor. It was noted that this analysis did not take into
account the change in reaction rate with temperature. However, the small temperature
changes per minute that occur in this analysis show that there would exist a large amount of
time to detect problems with the cooling coils via thermocouples installed in the reactor. Flow
meters installed in the pipes leading to the cooling coils and temperature could also detect
shutdown of the coils. To avoid this occurrence, multiple pumps should be installed in parallel
to feed the cooling coils with water. The adiabatic temperature rise calculations can be found in
Appendix I-D.
Table 7: Calculation of parameters for the simplified adiabatic methanol reactor energy balance
Computation of temperature of reactor versus time
Total volume
1094.027
m3
Volume of one reactor
364.6757
m3
Volume of oil
350.6497
m3
Density of oil
0.838
838
Total mass of oil
Heat capacity of oil
293844.4
1.67
g/mL
kg/m3
kg
kJ/kg-K
31
mCp of oil
490720.2
kJ/K
Mass flow rate of gas
49033.97
kg/hr
Heat capacity of gas
7.226894
kJ/kg-K
Residence time of gas
0.02
hr
mCp of gas
7087.266
kJ/K
Heat duty of reactor
6.61E+07
kJ/hr
2.20E+07
kJ/hr/reactor
Initial reactor temperature
210
483.15
°C
K
Simplified MeOH Reactor Temperature vs. Time
300
T (°C)
280
260
240
220
200
0
20
40
60
80
100
120
t (min)
Figure 5: Plot of reactor temperature (in °C) versus time (in min)
Life Cycle Analysis
A life cycle assessment (LCA) identifies, measures, and evaluates the environmental impact of
e e
stage of a p odu t s life a o di g to the i te atio al sta da ds set fo th ithi I“O
14040. These standards dictate the four stages to be analyzed in an LCA are: Goal and Scope
Definition, Inventory Analysis, and LCIA (life cycle impact assessment) and Interpretation.
Goal and Scope Definition
The goal of this LCA is to identify the environmental impacts of building and operating a solar
thermal biomass to methanol plant over the entirety of its operation, and then quantifying this
impact in terms of greenhouse gas (GHG) emissions, namely carbon dioxide, and other
influential factors. The LCA should be an integral part of the decision making process when
32
developing a new plant to ensure long term and short term public and land health and safety
issues are being addressed properly.
The plant operates with corn stover as the feedstock input, solar thermal, electricity, and
natural gas as the energy inputs, utilities as commodity inputs, and carbon dioxide, CO2,
emissions as the main output. An inventory analysis was performed on inputs and outputs for
pre-processing, gasification, and methanol purification, as well as a collective analysis over the
lifetime of the plant. Input data was generated from Aspen PLUS, the U.S Lifecycle Inventory
Database, and other resources. Figure 6 illustrates the system boundary used in this LCA.
The pre-processing unit inventory analysis involves analyzing the impact of the feedstock
agriculture, transportation, and CO2 absorption. By tracing the feedstock to its source, assuming
the corn is grown in Nevada, IA, the transportation emissions can be estimated. Land use,
fertilizer use and the resulting emissions were quantified as well and explained in detail in the
inventory analysis. Energy inputs to the plant include solar thermal energy, natural gas, and
electricity, with outputs being GHG emissions. Indirect effects, such as utility usage, are not
considered in this analysis.
33
Figure 6: System boundary for LCA for biomass to methanol plant
Inventory Analysis
The pre-processing unit uses corn stover as the biomass feedstock. The analysis of this product
assesses direct influences, such as CO2 emissions from transportation and CO2 absorption of
corn. As well, other factors were evaluated, such as the land use to grow the crop and fertilizer
needed for agriculture and the resulting emissions. A report from Iowa State University cites
the yield of corn stover to be 2.1 dry tons/acre (Zhang, Yanan, 2014). The biomass plant
requires an input of 82,000 tons/yr of feedstock, which equates to 39,048 acres of land/year to
34
provide enough corn stover to fuel the plant. When corn stover is used in biofuels production,
the entirety of the crop is depleted. Otherwise, the majority of the crop residue is left on the
fields to naturally replenish the soil with key elements such as nitrogen (N), phosphorous (P),
and potassium (K). Using all of the corn stover residue will require farmers to purposely
replenish the land with lost nutrients. Replacement rates for N, P, and K fertilizers were
calculated by Argonne Laboratory and are as follows: 7700 g N, 873 g P, and 9957 g K per 1 ton
of removed corn stover (Argonna National Laboratory, n.d.). The results of these calculations
are displayed in Table 8.
Table 8: Environmental factors from the waste feedstock
Factor
Amount
Unit
Land use
39,048
Acres land/year
Fertilizer added - Nitrogen
6.31x108
g N/year
Fertilizer added - Phosphorous
7.16x107
g P/year
Fertilizer added - Potassium
8.16x108
g K/year
N2O emission from added fertilizer
8.37x106
g N2O/year
NO emission from added fertilizer
4.10x108
g NO/year
Further, the CO2 emissions resulting in the transportation of the corn stover to the plant was
evaluated from Nevada, IA to Daggett, CA. Iowa is the largest producer of corn in the United
States, and thus feedstock originating in Iowa is a fair assumption. It was also assumed the
feedstock travels by a diesel powered train, as that is commonly used to transport freight
(Association of American Railroads, 2015). Train capacity varies substantially, anywhere from
1,000 to upwards of 20,000 tons. Frequency of travel would depend on the yield of corn stover
per crop per season, and how long corn stover can be stored after cultivation. BNSF reports 1
ton of freight is capable of travelling 423 miles on 1 gallon of diesel fuel (BNSF). After taking
into consideration an annual plant input of corn stover of 82,000 tons with a distance traveled
of ~1600 miles (from Nevada, IA to Daggett, CA), an approximate amount of CO2 emissions
from transportation fuel can be calculated. Michigan State University cites corn as having a
tremendous potential to remove carbon dioxide, upwards in the amount of 0.57 kg CO2/1 kg
35
corn stover (Michigan State University, 2007). The amount of CO2 absorbed by the required
amount of feedstock for the plant was assessed as well. The results of these calculations are
displayed in Table 9.
Table 9: CO2 emissions and absorbance due to transportation to plant and CO2 absorption
Feedstock
lb CO2/year
lb CO2/30 years (lifetime)
Transportation to plant
7.03x106
2.11x108
CO2 absorption from corn
9.33x107
2.8x109
The electrical energy and natural gas (if applicable) inputs for each section of the plant (preprocessing, gasification, and purification) along with subsequent CO2 emissions were assessed.
Natural gas is an input only in the gasification section to supply additional power to the hybrid
solar reactors. Using the assumptions given of 1.8 lb CO2/1 kWh electricity, 15 lb CO2/100 ft3
natural gas delivered, and 117 lb CO2/1 million BTU released on combustion, the CO2 emitted
from each section was calculated and summed. The CO2 absorbance of corn stover was
subtracted from this to estimate the total plant CO2 emissions. A comparison of total emissions
of the proposed hybrid biomass plant (type I) to a typical methanol production plant was
assessed. The Methanol Institute cites typical methanol plant emissions from fossil fuels as
1000kg CO2 per 2000lb methanol produced (Methanol Institute, n.d.). For an annual production
of 58.3 million gallons of methanol, this equates to 1.18x1010 lb CO2 over the lifetime of the
plant. This was also compared to a biomass plant with only solar inputs (type II). The results of
these calculations are in Table 10.
Table 10: CO2 emissions yearly and annually for various plant types
Type of Plant
lb CO2/year
lb CO2/30 years (lifetime)
Typical plant (fossil fuels)
3.93x108
1.18x1010
Biomass type I
2.16x108
3.68x109
Natural Gas
1.26x108
3.79x109
Electricity
8.97x107
2.69x109
CO2 Absorption
-9.33x107
-2.8x109
36
Biomass type II
-3.57x106
-1.07x108
Difference: typical vs biomass type I
1.77x108
8.09x109
Difference: Biomass type I vs type II
2.2x108
3.79x109
Impact Assessment and Interpretation
The inventory analysis confirms many factors need to be taken into consideration before using
waste biomass in a conversion to methanol plant. Farmers will have to address the issue of
fertilizer replacement, especially with low crop-yield seasons. This adds cost to their business
that may or may not be well-received. Interestingly, ash from the biomass gasification process
can be used as a fertilizer. This could present a possible solution for farmers. As well, modern
farming practices allow for more sustainable methods to retain soil nutrients, which might be
necessary if this becomes an issue with farmers.
The CO2 absorbance from the biomass more than outweighs the CO2 emittance from the
transportation of the feedstock to the plant, though this CO2 absorbance would likely rather be
used as a carbon credit in the plant economics. Thus, the transportation CO2 emissions has an
environmental impact that would need to be addressed. Utilizing larger capacity trains to travel
less often could potentially lessen the transportation emissions. As well, as fuel cell and battery
technology advances, converting the shipment to advanced technology, carbon free vehicles
would mitigate this issue entirely.
On comparison of CO2 emissions for a typical methanol plant and the proposed biomass hybrid
plant (type I), the latter cuts emissions by more than half. This is a substantial difference in GHG
emissions, though even more so if the reactor were to operate completely on solar. Biomass
type II plant would, theoretically, operate completely on solar without the need of natural gas,
perhaps through the use of thermal energy storage. Negating the use of natural gas alone
allows the plant to have negative carbon emissions, when including the CO2 absorbance in the
overall emissions calculation. Type I was also compared to type II, and the difference in
emissions was still on the order of 108. This calculation confirms the use of natural gas still
contributes substantially to the GHG emissions, though natural gas cuts CO2 emissions ~50%
versus coal (US Energy Information Administration, n.d.).
37
V. Project Premises
Included in the scope of this project are four systems working in tandem to produce high
quality methanol from waste biomass. The first system is a biomass pre-processing section
modeled in SuperPro Designer which dries, grinds, and pressurizes corn stover biomass. The
second system is modeled in Aspen PLUS. The biomass is converted to syngas, which passes
through many clean-up steps and is converted into methanol. The third system is the modelling
of one of these clean-up steps in Aspen HYSYS: an amine scrubbing system which removes H2S
and CO2 from the syngas. The fourth system is a post-processing operation modeled in Aspen
HYSYS to purify the methanol product. These systems combined constitute the methanol
production plant.
The methanol production plant was designed to continuously produce 58,300,000 US gallons of
99.97 mol% of methanol per year. Slight excess of methanol was produced to account for
unforeseen major maintenance issues that could cut into production time. In this manner, the
u e s suppl
ould ha e li ited i te uptio s o e the ou se of the pla t s lifeti e, if a
at
all. It was assumed that all raw materials are available at 100% purity. This plant was designed
to be a new g ass oots pla t, so the osts of la d, site p epa atio , royalties, and other
related expenses were considered. The plant was planned to be constructed and operate in
Daggett, California, which is located in unincorporated San Bernardino County. It was assumed
that all utilities, such as cooling water, steam, electricity, wastewater systems, and compressed
air are readily available from nearby industrial sites. Other assumptions and specifications
unique to each system are outlined below.
Design
Biomass Pre-Processing




Access to preheated dry air at 120°C and 1.013 bar assumed
Access to nitrogen gas at 25°C and 35 bar assumed
Dryer provides an evaporation rate of 20 (kg/hr)/m3
Power Requirements:
o Shredder: 47 kW/(kg feed/s) with 0% power dissipation to heat
38
o Hammermill grinder: 130 kW/(kg feed/s)
o Lockhopper: 490 kW/m3, adiabatic operation, 5% nitrogen leaks to gasifier and
the remainder is vented by the hopper
o The required particle size distribution is met with the given power inputs
Biomass Gasification and Methanol Production











Fluids package: RK-Soave with unconventional components
Processed biomass enters the system at 90°C and 35 bar
Gasification reactor modeled as decomposition of biomass into elemental forms fed into
an equilibrium (Gibbs) reactor
1/3 of the duty of gasification provided by concentrated solar power, 2/3 provided by
natural gas burners
Conversion of biomass in gasification reactor assumed to be 100%
o Kinetics modeled in Excel
Conversion of H2S and HCl in ZnO reactor assumed to be 100%
o Kinetics modeled in Excel
Conversion of CO and CO2 in methanol reactor assumed to be 45%
o Kinetics modeled in Polymath
Catalyst in methanol reactor is Cu/ZnO/Al2O3 in slurry phase with Witco-40 mineral oil
Efficiency of solid separations (cyclones) assumed to be 100%
Pressure drops through piping and equipment assumed to be negligible
All water and natural gas streams fed at 25°C
Amine Scrubbing


Fluids package: Acid Gas Cleaning
Specified to reduce the concentration of H2S in sweetened syngas to <1 ppm
Methanol Purification


Fluids package: Peng-Robinson
Final product volume calculated at a methanol density of 735 kg/m3 at 80°C (The
Engineering Toolbox, n.d.)
39
Economics




Bare module cost method used to determine project capital cost
o Whe
i stalled used i assu ptio s he e, a e
odule fa to is
Base cost of heliostats is $126/m2, installed
Base cost of secondary concentrator mirror area estimated at 10X heliostat cost per m 2
($1260/m2), installed
Base cost of solar tower (installed) in 2015 given by Equation (21) as follows, where ℎ is
the height in m:

















= .
,
+
.
.
(21)
Cost of natural gas: $2/1000 SCF (standard cubic feet)
Cost of biomass: $60/metric ton delivered by railcar on a dry basis
Cost of ZnO and methanol catalyst: $5/kg catalyst
Value/cost of 450 psig (high pressure) steam: $17.29/1000 kg
Value/cost of 150 psig (medium pressure) steam: $12.57/1000 kg
Value/cost of 50 psig (low pressure) steam: $7.86/1000 kg
Price of cooling water: $0.019/m3
Refrigeration, -150°F: $15/GJ
Refrigeration, -90°F: $12.21/GJ
Refrigeration, -30°F: $9.43/GJ
Refrigeration, 10°F: $6.57/GJ
Chilled water, 40°F: $4.71/GJ
Wastewater and solid hazardous waste treatment: $0.31/kg contaminant removed
Landfill: $0.19/dry kg
Cost of land: $1000/acre
SiC tubes for the solar reactor are 6 inches in I.D., ¾ inch thick, and up to 20 m long
SiC tubes costed as follows (email communication with Prof. Weimer):
o First reactor tube costs $1M, each other tube after first costs $300,000
40













Cost of alumina insulation, metal containment shell, and secondary concentrator are all
installed
5 yr MACR depreciation
Cost of borrowed money (interest on capital) is 10%
Operation is 24 hr/day for 333 days/yr (8000 hr/yr)
Plant capacity is 50% in year 1, 75% in year 2, and 100% thereafter
Lifetime of the plant is 30 years
Construction period of the plant is 1 year
Contingency is 15%
Inflation is 1.9% throughout plant life for all considerations
Effective tax rate is 38.9%
Insurance and local taxes is 2%
Total fixed cost used in cash flow calculations
Cost of labor (annual wages per operator) is $104k/operator/shift (includes all
overhead)
41
VI. Approach
Hand Calculations
Heat of Reactions: Cellulose and Steam
Heat of Reaction at 25°C
The first hand calculation conducted in the approach calculations determines the heat of
reaction (∆
), when cellulose is reacted with steam.
In order to obtain this value, the following reaction was provided:
+
→
+
(22)
), which is −
This equation has a known heat of combustion (∆
.
/ at 25°C.
Using this equation, the enthalpy of cellulose can be back calculated. The heat of combustion is
the difference between the enthalpies of the products and the enthalpies of the reactants
multiplied by their stoichiometric coefficients:
∆
=[
+
�
�
]−[
+
�
]
(23)
Here, the enthalpies of carbon dioxide, water and oxygen gas at 25°C can be found in species
reference tables, such as those in Pe y’s Che i al E gi ee i g Ha d ook. With these values,
the only unknown is the enthalpy of cellulose,
.
cellulose was determined to be
/
. By this method, the enthalpy of
.
In order to determine the heat of reaction of cellulose when reacted with steam, the chemical
formula for this desired reaction is given:
+
→
+
(24)
Using the enthalpy of cellulose, the heat of reaction at 25°C can be found by once again taking
the difference between the products and reactants:
∆
=[
�
+
�
]−[
+
By this method, the heat of reaction at 25°C was determined to be −
�
]
.
(25)
/
.
42
Heat of Reaction at 1450°C
In order to calculate the heat of reaction at high temperature, the enthalpies of the reactants
and products involved in the reaction must first be raised to their values at that temperature.
This is achieved by integrating the constant pressure heat capacities of each product and
reagent across the observed temperature increase:
=∫
(26)
After this integration has been performed on each product and reagent, Equation (25) is used
again, and the heat of reaction is obtained. At 1450°C, the heat of reaction for the
decomposition of cellulose in the presence of steam was found to be −
.
/
.
Detailed analyses of these calculations can be found in Appendix I-A: Approach Calculations.
Heat of Reaction: Lignin and Steam
The determination of the heat of reaction of lignin in the presence of steam at 25°C and 1450°C
was accomplished by the same method used to find the heat of reaction of cellulose in the
presence of steam. By these calculations, it was determined that the decomposition of lignin in
the presence of steam at 25°C had a heat of reaction of −
reaction at 1450°C had a heat of reaction of −
.
/
.
/
, while the same
. A detailed description of the
calculations performed in order to achieve these values is available in Appendix I-A: Approach
Calculations.
Waste Biomass Feed Estimation
In order to estimate the amount of biomass required as system feed, the number of moles of
biomass produced on a per hour basis was divided by the 45% conversion of the methanol
reactor while ignoring the recycle streams. Knowing that one mole of CO was produced per
mole of biomass reacted, and that one mole of methanol was produced per mole of CO reacted
in the methanol reactor led to an initial guess of the moles of CO required in total. The project
description stated the optimal molar ratio of H2 to CO was 2, and that the steam methane
reforming reaction would lead to a 3:1 mole ratio of H2:CO per mole of methane reacted. The
43
final piece of given information was the weight percent of the three components fed to the
reactor, as shown in Table 11.
Table 11. Weight percent of cellulose, lignin, and ash in the feed gas.
wt% Cellulose
73.36%
wt% Lignin
23.34%
wt% Ash
1.18%
The first step in the calculation was to convert the percent mass composition to a mole percent
composition. This was done by taking a basis of 100g and using the molecular weights of each
component to determine the number of moles that would be present, then dividing each
component by the total number of moles. The results of this calculation are summarized in
Table 12.
Table 12. Mole percent results from hand calculation.
mol% Cellulose
75.21%
mol% Lignin
21.53%
mol% Ash
3.26%
Once the mole percent of the feed biomass was determined, stoichiometry was used to
determine the H2:CO ratio produced by biomass gasification. The results of this calculation were
that 1.067 moles of H2 were produced per mole of CO that was produced. Knowing that 2 moles
of Hydrogen per mole of carbon monoxide was the goal, an iteration was set up to determine
the fraction of CO produced by biomass, knowing that the fraction of CO produced by hydrogen
would be the difference between 1 and the fraction of CO produced by biomass. This iteration
was checked by ensuring that the overall moles of H2 produced was twice that of the overall
moles of CO produced.
Once the mole fraction of fed methane and fed biomass was determined, the total moles of CO
produced was used to calculate the feed rates of biomass and methane to the reactor using a
one-to-one ratio of moles CO produced to moles biomass or methane reacted.
44
Theoretical Energy Requirement for Solar-Thermal Reactor
The final step in the approach calculations is to determine the theoretical energy requirement
for the solar-thermal reactor. In order to do this, the power required to raise the temperatures
of the biomass, methane and steam to 1450°C is calculated.
These calculations are accomplished by integrating the constant pressure heat capacities of
each component across the desired temperature change, as shown in Equation (26). For lignin
and cellulose, the change in enthalpy for each component is added to ∆
for each reaction,
respectively. This value is then multiplied by the given molar flowrate of each component in
order to obtain the power required to heat, and recovered from reacting, a given biomass
substrate. This power can then be converted to kW.
This procedure is repeated independently for lignin, cellulose, steam and methane. The
individual power requirements for each component as well as the total estimated power
provided by the reactor are provided in the following table. Detailed calculations used to obtain
these values are present in Appendix I-A: Approach Calculations.
Table 13. Summary of power input/output for solar-thermal reactor.
Component
Power Input / Output (+/-)
Cellulose
-54046 kW
Lignin
-23432.7 kW
Steam
17131.8 kW
Methane
25341.6 kW
Total:
-35005.5 kW
45
VII. Process Flow Diagrams with Material & Energy Balances
Below are the process flow diagrams (PFD) with associated material and energy balances for
the pre-processing, gasification, amine scrubbing, and post-processing sections. Detailed
material and energy balance calculations can be found in Appendix I-B. Images for all of the
simulations may be found in Appendix III-A through Appendix III-D.
Biomass Pre-Processing
Process Description and PFD
Shown below in Figure 7 is the process flow diagram for the biomass pre-processing simulation.
This section is the first stage of the entire process, where the corn stover bales are made ready
for the gasification reactions that will occur further in the system.
Figure 7: Pre-processing section PFD
46
The feed enters at point S-101 at room temperature and pressure, where it is conveyed on a
rolling belt system to the first stage of pre-processing. The process begins with 2.55 x 104
corn stover with a composition shown below in Table 14.
ℎ
of
Table 14: Composition of Biomass Mixture with Pure Components
Component Name
Mass %
1
Ash
0.8700
2
Carbon
38.3920
3
Chlorine
0.0380
4
Hydrogen
4.5600
5
Nitrogen
0.1500
6
Oxygen
30.9750
7
Sulfur
0.0150
8
Water
25.0000
The stage 1 shredder will do the initial size reduction such that the corn stover is no larger than
� . in diameter. This is necessary before entering the dryer as trying to dry an uncut bale is
too time- and energy-intensive.
The material is then moved via rolling conveyer to a direct-air rotary drum dryer system that
will reduce the moisture content from 25% to 6.25%. This operation is done to prevent excess
water from entering the solar reactor as well as to reduce complications in the later grinding
stage where the excess moisture can clump and block the grinder and lock-hopper. The 2.55 ×
104
ℎ
of air enters at point S-102 at 120°C and exits at point S-104 at 90°C.
The remaining biomass enters the stage 2 grinder, P-3 / GR-101 Grinding, which was designed
as a hammer mill. This unit reduces the size of the particles to that necessary for the
gasification reaction as well as pressurization in the lock-hopper and the size distribution is
shown below in Table 15.
47
Table 15: Biomass Particle Size Distribution from Hammer Mill.
Size Interval Lower Limit (�
Size Interval Upper Limit (� )
Weight Fraction Biomass Particles
100
120
0
120
140
.1
140
160
.2
160
180
.3
180
200
.4
The biomass then flows via pneumatic conveyance to the lock-hopper. Pneumatic conveyance
was used to prevent particles of this size to be lost from the system. The lock-hopper, P-4 / HP101, is capable of pressurizing the biomass to 35.0 bar, which is necessary to achieve the partial
pressures inside the solar reactor. This is achieved by adding 2.55 x 103
ℎ
nitrogen at 35.0 bar
and 25°C in the bottom section of the lock-hopper. 95% of the nitrogen is bled from the system
and 5% of the nitrogen leaks into the biomass being fed to the solar reactor. This was modeled
in the PFD by P-5 / CSP-101 Component Splitting. After this step, the biomass is successfully
dried, sized, and pressurized as required by the system. The final feed to the solar reactor
leaves via S-110 at 85.59°C and 35.0 bar with a flowrate of 2.05 x 104
Material Balances
ℎ
.
An atomic material balance was carried out on the biomass pre-processing section to ensure
closure of the system. Using Equation (27) below where A is some atomic species and there is
no generation, Table 16 was generated.
∑ ̇
,�
−∑ ̇
,
+
�
=
�
(27)
Table 16: Atomic balance around the biomass pre-processing section
Atomic Species
In
Out
Accumulation (lbmol/hr)
(lbmol/hr)
(lbmol/hr)
Ash
1.11111
1.11111
0
Carbon
816.4497
816.4497
0
Chlorine
0.13689
0.13689
0
Hydrogen
1286.683
1286.683
0
48
Nitrogen
1351.599
1351.599
-0.00042
Oxygen
936.3302
936.3298
.000426
Sulfur
0.11949
0.11949
0
From the energy balance, the system appears to be closed, and the only slight deviation is in
nitrogen and oxygen, which is most likely due to the number of significant digits used in the
mol% calculation in air. The error was insignificant, and the system is considered closed.
Heat Duty
Energy inputs and outputs were analyzed for the biomass pre-processing section. All the
required duties are supplied via electricity; a more in-depth analysis of the utility cost is
included in Section IX. Shown below in Table 17 is a summary of the heat duty per unit, where
positive duties represent the addition of energy to the unit operation.
Table 17: Heat duty summary for Biomass Pre-Processing Section
Unit Operation
P-1 / SR-101 Shredding
P-2 / RDR-101 Rotary Drying
P-3 / GR-101 Grinding
P-4 / HP-101 Hopper
Heat Duty Required (Btu/hr)
Provided by:
.
Electricity
.
Electricity
.
Electricity
.
Electricity
Biomass Gasification
Process Description and PFD
The following figure illustrates the primary section of the Solar-Thermal Biomass Gasification
facility, in which the biomass is converted to methanol. An explanation of the unit operations
and material flow within them, including flow rates, temperature, and pressure, is explained
below. A PFD for this section of the plant is shown in Figure 8.
49
Figure 8: Biomass gasification PFD
50
Biomass from the pre-processing section enters the solar reactor. In Aspen PLUS, the solar
reactor was modeled as two separate units. The first unit (DECOMP) was used to model the
breakdown calculations, while the second unit (SOLAR) models a Gibbs equilibrium reactor.
FORTRAN code was used with DECOMP in a breakdown calculator (BRKDOWN) to perform this
function; see Appendix III-F for this code. The biomass stream enters the DECOMP breakdown
unit at 90.0°C and 35.0 bar with a mass flowrate of 2.05×104 lb/hr. The heat duty of the
breakdown reactor is 7.75×107 Btu/hr. After undergoing breakdown and exiting the unit, the
stream temperature is raised to 1450°C with no change in pressure. In the SOLAR reactor, the
biomass process stream is contacted with 1.13×104 lb/hr of methane at 25.0°C and 35.0 bar. A
pressurized water stream (H2O-SLR) also enters the SOLAR reactor with a flowrate of 1.92×104
lb/hr at 26.0°C and 35.0 bar. The final stream that enters the solar reactor is a recycle stream
(R-SOLAR) from a later portion of the methanol synthesis process. This recycle stream has a
flowrate of 6.28×103 lb/hr and enters the solar reactor at a temperature of 50.0°C and 35.0 bar.
The heat duty of the solar reactor is 1.45×108 Btu/hr. Both this heat duty and the heat duty
supplied to the DECOMP breakdown unit are provided by the solar field and natural gas
burners. In the breakdown and solar reactor, the biomass is broken down into hydrocarbons
and other small molecules. A material balance associated with this reactor is given in Table 18.
Table 18: Material balance around the solar reactor (DECOMP + SOLAR)
IN
Stream:
BIOMASS
R-SOLAR
OUT
H2O-SLR
METHANE
PRODUCTS
Component Mass Flow Rates (lb/hr)
N2
0
338.352
0
0
514.3805
C2H6
0
0.005
0
0
0.009
CH3OH
0
255.106
0
0
0
C2H4
0
0.115
0
0
0.193
HCl
0
0
0
0
13.841
H2O
0
9.098
19200
0
1890.892
H2S
0
0
0
0
6.131
H2
0
709.237
0
0
6792.596
0
6.92×10
0
0
1.05×10-5
NO
6
51
NH3
0
1.99E-07
0
0
3.383
CO
0
4850.051
0
0
46935.05
CO2
0
1.825
0
0
665.592
CH4
0
119.140
0
11250
183.983
NO2
0
0
0
0
7.84×10-14
Ash
0
0
0
0
226.878
Biomass
20500
0
0
0
0
TOTAL
20500
6282.929
19200
11250
57232.93
TOTAL IN
(lb/hr)
TOTAL OUT
(lb/hr)
DIFFERENCE (lb/hr)
57232.93
57232.93
0.00
The product stream from the solar reactor exits at 1450°C and 35.0 bar. The flowrate of this
stream is 5.72×104 lb/hr. In order to cool this stream, water is fed into a quench tank (SPRAY-Q)
where it is contacted with pressurized cooling water (H2O-IN). The amount of cooling water
required to achieve the desired process stream temperature reduction is 5.95×104 lb/hr. The
cooling water enters at 25.5°C and 35.0 bar, and exits at 586°C. The cooling duty of the quench
tank is 9.49×107 Btu/hr, and the process stream is cooled to 120°C with no pressure change.
This process stream is then mixed with 2.98×102 lb/hr of the cooling water stream exiting the
quench tank. Contact with the 586°C cooling water stream raises the process stream
temperature to 122.0°C. A material balance for the quench tank system including the mixing
point between the exiting utility stream and process stream is given in Table 19.
Table 19: Material balance around the quench tank
IN
Stream:
PRODUCTS
OUT
H2O-2
QNCH-H2O
TO-CYCL
Component Mass Flow Rates (lb/hr)
N2
514.3805
0
0
514.3805
C2H6
0.009
0
0
0.009
C2H4
0.193
0
0
0.193
HCl
13.842
0
0
13.842
H2O
1890.892
59524.74
59227.12
2188.516
H2S
6.131
0
0
6.131
52
H2
6792.596
0
0
6792.596
NO
1.05×10
0
0
1.05×10-5
NH3
3.383
0
0
3.383
CO
46935.05
0
0
46935.05
CO2
665.592
0
0
665.592
CH4
183.983
0
0
183.983
NO2
7.84×10-14
0
0
7.84×10-14
Ash
226.878
0
0
226.878
TOTAL
57232.93
59524.74
59227.12
57530.55
-5
IN (lb/hr)
OUT (lb/hr)
DIFFERENCE (lb/hr)
116757.67
116757.67
0.00
The syngas stream exiting the quench tank system then enters a cyclone (CYCLONE), where
2.27×102 lb/hr of solid wastes are removed from the process. A material balance for the cyclone
separation unit is given in Table 20.
Table 20: Material balance around the cyclone (CYCLONE)
IN
Stream:
TO-CYCL
OUT
VAPOR
SLDWASTE
Component Mass Flow Rates (lb/hr)
N2
514.381
514.381
0
C2H6
0.009
0.009
0
C2H4
0.193
0.193
0
HCl
13.842
13.842
0
H2O
2188.516
2188.516
0
H2S
6.131
6.131
0
H2
6792.596
6792.596
0
NO
1.05×10-5
1.05×10-5
0
NH3
3.383
3.383
0
CO
46935.05
46935.05
0
CO2
665.592
665.592
0
CH4
183.983
183.983
0
Ash
226.878
0
226.878
TOTAL
57530.55
57303.67
226.878
53
IN (lb/hr)
OUT (lb/hr)
DIFFERENCE (lb/hr)
57530.55
57530.55
0.00
The remaining vapor exiting the cyclone is further cooled through a heat exchanger (NH3-SEP
HX) that uses 2.99×105 lb/hr of refrigerated brine, entering at -17.8°C and atmospheric
pressure. The refrigerated brine performs 9.16x105 Btu/hr of cooling duty on the process
stream, which exits the cooler at 20.0°C and 5.73×104 lb/hr. This stream then enters the NH3SEP unit, in which aqueous waste is separated from the process. This aqueous waste consists
primarily of H2O and NH3. A complete material balance for the NH3-SEP unit is provided in Table
21.
Table 21: Material balance around NH3-SEP
IN
Stream:
VAPOR
OUT
AQ-WASTE
ACID-GAS
Component Mass Flow Rates (lb/hr)
N2
514.381
1.82×10-5
514.381
C2H6
0.009
3.64×10-9
0.009
C2H4
0.193
1.20×10-7
0.193
HCl
13.842
1.99×10-3
13.840
H2O
2188.516
2139.523
48.992
H2S
6.131
0.003
6.128
H2
6792.596
4.18×10-4
6792.596
NO
1.05×10-5
8.72×10-12
1.05×10-5
NH3
3.383
1.802
1.581
CO
46935.05
0.001
46935.05
CO2
665.592
0.009
665.582
CH4
183.983
6.78×10-5
183.983
TOTAL
57303.67
2141.341
55162.33
IN (lb/hr)
OUT (lb/hr)
DIFFERENCE (lb/hr)
57303.67
57303.67
0.00
54
The remaining acid gas stream exiting the NH3-SEP unit then enters the AG-CLEAN unit, which
represents the amine scrubbing section. This section is described in detail later in this report. A
summary material balance for this section is provided in Table 22 below. This material balance
also includes the split fractions (seen in Table 23) achieved in the acid-gas cleaning process,
where the split fractions are on a molar basis and represent the percentage of the two inlet
streams which leave the CLN-GAS stream. The stream XTRA-H2O represents the additional
water which vaporizes in the amine scrubbing system. The material balance is not closed to the
third decimal place because the material balance was solved in Aspen PLUS according to the
low tolerance of the numerical solver. Detailed calculations for the split fraction can be found in
Appendix I-D.
Table 22: Material balance around AG-CLEAN
IN
Stream:
ACID-GAS
OUT
XTRA-H2O
CLN-GAS
ACIDS
Component Mass Flow Rates (lb/hr)
N2
514.381
0
514.362
0.018
C2H6
0.009
0
0.009
1.93×10-6
C2H4
0.193
0
0.193
1.06×10-4
HCl
13.840
0
0
13.840
H2O
48.992
49.851
98.843
0
H2S
6.128
0
0.118
6.010
H2
6792.596
0
6792.324
0.272
NO
1.05×10-5
0
1.05×10-5
0
NH3
1.581
0
1.66×10-6
1.581
CO
46935.050
0
46932.7
2.347
CO2
665.582
0
174.249
491.333
CH4
183.983
0
183.965
0.017
TOTAL
55162.33
49.851
54696.77
515.417
IN (lb/hr)
OUT (lb/hr)
DIFFERENCE (lb/hr)
55212.18
55212.1869
-0.006
55
Table 23: Molar split fraction from HYSYS for AG-CLEAN
Component
Molar Split
Fraction
N2
1
C2H6
1
C2H4
1
HCl
0
H2O
-
H2S
0.019
H2
1
NH3
0
CO
1
CO2
0.262
CH4
1
Clean gas exiting the amine scrubbing subsystem is stripped of H2S, HCl and NH3. This
procedure leaves a stream primarily composed of N2, H2O, CO, CO2 and CH4. This clean gas
stream is then heated by a high-pressure steam shell-and-tube heat exchanger (HEAT-1), which
performs 1.23×107 Btu/hr duty on the process stream. The high pressure steam stream used to
achieve the desired temperature change enters the exchanger at a mass flowrate of 1.61×104
lb/hr, 231°C and 31 bar. This stream exits the exchanger at the same temperature and pressure,
having only undergone condensation when imparting energy to the process stream. The
process stream exits the exchanger at 210°C and 35.0 bar prior to entering the zinc reactor (ZNREACT). The zinc reactor removes any remaining H2S from the process stream using a zinc
catalyst. The material balance surrounding the zinc reactor is detailed in Table 24.
Table 24: Material balance around ZN-SPLIT
IN
Stream:
TO-ZN-R
OUT
ZNO
ZN-PROD
Component Mass Flow Rates (lb/hr)
N2
514.362
0
514.362
56
C2H6
0.009
0
0.009
C2H4
0.193
0
0.193
H2O
98.843
0
98.905
H2S
0.118
0
0
H2
6792.324
0
6792.324
NO
1.05×10-5
0
1.05×10-5
NH3
1.66×10-6
0
1.66×10-6
CO
46932.700
0
46932.700
CO2
174.249
0
174.249
CH4
183.965
0
183.965
ZnO
0
0.281
ZnS
0
0
1.84×10-6
0.337
TOTAL
54696.77
0.281
54697.05
IN (lb/hr)
OUT (lb/hr)
DIFFERENCE (lb/hr)
55212.18
55212.1869
-0.006
The process stream exits the zinc reactor at 210°C and 35.0 bar. This stream then enters a
second cyclone (ZN-SPLIT) in which 3.37×10-1 lb/hr of zinc-containing solids are removed from
the process. The remaining vapor stream, (ZN-VAP) is contacted with 5.34×104 lb/hr of a recycle
stream. This recycle syngas stream contains primarily hydrocarbons and hydrocarbon
compounds including C2H6, H2, CO, CO2, CH4 and CH4O as well as N2 and H2O. A material balance
around this mixing point is given in Table 25 below.
Table 25: Material balance around MIX-2
IN
Stream:
ZN-VAP
OUT
RECYCLE
TO-COMP
Component Mass Flow Rates (lb/hr)
N2
514.362
2875.990
3390.269
C2H6
0.009
0.046
0.055
CH3OH
0
2168.401
2168.377
C2H4
0.193
0.981
1.174
H2O
98.905
77.329
176.233
H2
6792.324
6028.513
12820.720
57
NO
1.05×10-5
5.88×10-5
6.93×10-5
NH3
1.66×10-6
1.69×10-6
3.35×10-6
CO
46932.700
41225.430
88158.130
CO2
174.249
15.512
189.761
CH4
183.965
1012.694
1196.655
TOTAL
54696.71
53404.9
108101
IN (lb/hr)
OUT (lb/hr)
DIFFERENCE (lb/hr)
108101.61
108101
0.61
The mixed process stream has a mass flowrate of 1.08×105 lb/hr, and enters a centrifugal
compressor at 133°C and 35 bar. The compressor performs 1.88×107 Btu/hr duty on the stream,
raising the pressure to 80.0 bar and the temperature to 279°C. This highly pressurized stream
then enters the final methanol reactor (REACT).
The methanol reactor is equipped to remove 6.26×107 Btu/hr of heat from exothermic
reactions which produce methanol. A pressurized water stream with a flowrate of 5.6×103
lb/hr also enters the process unit at 80 bar to provide the desired 10:1 ratio of CO:H2O. The
methanol reactor converts the CO, CO2, H2O, and H2 in the process stream to methanol with
impurities and water. A material balance surrounding this reactor is presented in Table 26
below. Detailed descriptions of the methanol reactor equations are given under the design
description of this unit.
Table 26: Material balance around REACT
IN
Stream:
TO-RCT
OUT
H2O-RCT
RCT-PROD
Component Mass Flow Rates (lb/hr)
N2
3390.269
0
3390.269
C2H6
0.055
0
0.055
CH3OH
2168.377
0
47611.890
C2H4
1.174
0
1.174
H2O
176.233
5600
5846.144
H2
12820.720
0
7094.891
NO
6.929×10-5
0
6.929×10-5
58
NH3
3.349×10-6
0
3.349×10-6
CO
88158.130
0
48541.320
CO2
189.761
0
18.976
CH4
1196.655
0
1196.655
TOTAL
108101
5600
113701
IN (lb/hr)
OUT (lb/hr)
DIFFERENCE (lb/hr)
113701.00
113701
0.00
The products exiting the methanol reactor are then cooled by a heat exchanger (VL-SEP HX)
prior to entering a vapor-liquid separator (VL-SEP). This heat exchanger removes 4.59×107
Btu/hr of heat from the methanol product stream, decreasing its temperature from 210°C to
50°C prior to entry into the VL-SEP unit. 1.53×106 lb/hr of cooling water is required to achieve
the desired temperature change. The cooling water enters the heat exchanger at 32.2°C and
atmospheric pressure and exits at 48.9°C.
The product-containing process stream then proceeds to a vapor-liquid separation unit (VL-SEP)
in which a raw methanol product stream exits the system, while the remaining vapor exits as
the SEP-VAP stream. The raw methanol product is sent to a methanol purification system,
which was modelled in HYSYS. A material balance surrounding the VL-SEP unit is provided
below in Table 27.
Table 27: Material balance around VL-SEP
IN
Stream:
OUT
RCT-PROD
RAW-MEOH
SEP-VAP
Component Mass Flow Rates (lb/hr)
N2
3390.269
6.751
3383.518
C2H6
0.055
0.001
0.054
CH3OH
47611.890
45060.830
2551.060
C2H4
1.174
0.020
1.154
H2O
5846.144
5755.168
90.975
H2
7094.891
2.523
7092.368
NO
NH3
6.929×10
-7
1.215×10
6.916×10-5
3.349×10-6
1.360×10-6
1.989×10-6
-5
59
CO
48541.320
40.810
48500.510
CO2
18.976
0.727
18.249
CH4
1196.655
5.251
1191.404
TOTAL
113701
50872.08
62829.29
IN (lb/hr)
OUT (lb/hr)
DIFFERENCE (lb/hr)
113701.00
113701.37
-0.37
The SEP-VAP stream is then split into three fractions. Two fractions of this stream form the
recycles that are fed the solar reactor (15%) and mix with the ZN-VAP stream (80%),
respectively, while the third fraction exits the system as a purge (5%). All streams exit at 50.0°C
and 35.0 bar. The solar reactor feed recycle stream has a mass flowrate of 6.28×103 lb/hr, while
the recycle that mixes prior to entering the compressor has a mass flowrate of 5.34×104 lb/hr.
The purge stream has a flowrate of 3.14×103 lb/hr. A final material balance surrounding the
splitting point is given in Table 28.
Table 28: Material balance around SPLIT-2
IN
Stream:
OUT
SEP-VAP
PURGE
RECYCLE
R-SOLAR
Component Mass Flow Rates (lb/hr)
N2
3383.518
169.176
2875.990
338.352
C2H6
0.054
0.003
0.046
0.005
CH3OH
2551.060
127.553
2168.401
255.106
C2H4
1.154
0.058
0.981
0.115
H2O
90.975
4.549
77.329
9.098
H2
7092.368
354.618
6028.513
709.237
NO
6.916×10-6
3.458×10-6
5.879×10-5
6.916×10-6
NH3
1.989×10-6
9.943×10-8
1.690×10-6
1.989×10-7
CO
48500.510
2425.025
41225.430
4850.051
CO2
18.249
0.912
15.512
1.825
CH4
1191.404
59.570
1012.694
119.140
TOTAL
62829.29
3141.465
53404.9
6282.929
60
IN (lb/hr)
OUT (lb/hr)
DIFFERENCE (lb/hr)
62829.29
62829.294
0.00
An overall material balance was then done around the whole system. The result of this material
balance is shown in the table below.
IN (lb/hr)
OUT (lb/hr)
DIFFERENCE (lb/hr)
116124.87
116124.638
0.23
Heat Duty
The heat duties for each piece of process equipment in the biomass gasification subsystem are
provided below, where a positive value is an input and a negative value is an output.
Table 29. Unit operation and heat duty summary for biomass gasification subsystem.
Unit Operation
Heat Duty or Power
Provided by:
Required (Btu/hr)
PUMP-1/Solar Reactor Pump
7.04×104
Electricity
DECOMP/Solar Reactor
7.75×107
Solar Field
SOLAR/Solar Reactor
1.45×108
Solar Field, Natural Gas
SPRAY-Q/Spray Quench Tank
-9.49×107
Cooling Water
PUMP-2/Pump for Spray Quench Tank
1.57×105
Electricity
NH3-SEP HX/Amine Separator Heat Exchanger
-9.16×106
Refrigerated Brine
HEAT1
1.23×107
High Pressure Steam
COMPRESS
1.88×107
Electricity
PUMP-3
6.47×105
Electricity
REACT
6.26×107
Electricity
VL-SEP HX
-4.59×107
Cooling Water
61
Amine Scrubbing
Process Description and PFD
Figure 9: Amine scrubbing process flow diagram
62
Following the gasification of the biomass and removal of a portion of the ammonia from the
process stream, it was desired to remove additional H2S and CO2 contaminants from the syngas.
H2S and CO2 are common contaminants in the conversion of both petroleum-based liquid fuels
and biomass-based liquid fuels, and must be removed for a few reasons. First, these gases
lower the activity of the catalyst in the methanol reactor downstream by poisoning it, lowering
the ea to s o e all o e sio effi ie
a d the lifeti e of the atal st (Seo, 2013). Second,
these gases form acidic compounds upon exposure to atmospheric moisture. H2“ fo
eak fo
of sulfu i a id, a d CO2 fo
s a
s a o i a id; the efo e, these gases a e ofte
efe ed to as a id gases a d should e e o ed to p e e t the o osio of e uip e t
(Arnold & Stewart, 1999). The process of removing sulfur acid gas components is known as
s eete i g. Thi d, H2S and CO2 reduce the heating value of the final product, and H2S can
release sulfur oxide (SOx) emissions upon combustion (Arnold & Stewart, 1999).
Amine scrubbing systems are frequently used in the treating of natural gas, and there exists
many approaches to remove acid gases from these fuels. Many of these processes involve the
use of aqueous amine solutions. The underlying chemistry involves the reaction of the basic
amine molecule to the contaminant acid gas to form a salt, which dissolves in water and is
thereby removed from the gas phase. Equations (28) and (29) provide for a simple set of
chemical equations for this overall chemistry (Beychok, 2012):
where
+
+
+
↔
↔
+
+
+
+
−
(28)
−
(29)
represents moieties attached to the central amine. These equilibria favor the products
at low temperatures. Primary, secondary, or tertiary amines can be used in acid gas treating; in
the proposed process, methyldiethanolamine (MDEA) was used because it is a tertiary amine,
and so is less basic and corrosive than its primary and secondary counterparts. MDEA can also
effectively bind H2S at high concentrations. When the amine must be regenerated by heating to
push the equilibria in Equations (28) and (29) to the reactants, MDEA can release its acid gases
with less heating duty per mole than other amines (Amines & Plasticizers Limited, n.d.). Finally,
MDEA is environmentally inert and biodegradable (Union Carbide Corporation, 2015). When
63
the MDEA solutio
solutio
o tai s o a id gas o po e ts, it is efe ed to as lea a i e. The
e o es i h o e a id gas o po e ts a e p ese t.
To simulate this amine chemistry for the proposed process, an amine scrubbing system was
implemented and simulated using the Acid Gas Cleaning simulation environment in Aspen
HYSYS (Aspen Technology, Inc., 2013). This simulation package is more effective at modeling
the complex acid-base and electrolyte chemistry involved in amine scrubbing than Aspen PLUS;
therefore, the HYSYS simulation is developed separately from the Aspen PLUS simulation. The
process and equipment to treat the syngas and regenerate and makeup the MDEA solution is
described in the following sections. A PFD of the process is shown in Figure 9. A screenshot of
the Aspen HYSYS simulation is shown in Appendix III-C.
To remove the acid gas contaminants, the process syngas (Feed Gas) is bubbled up through an
absorption column and flows countercurrently to the lean amine solution (Lean Amine). The
rich amine stream (Rich Amine) leaves the bottom of the column and the sweet gas stream
(Sweet Gas) leaves the top of the column. The absorber was assumed to operate adiabatically.
In accordance with the design specifications of the amine scrubbing system, the sweet gas
leaves the process with a concentration of H2S less than 1 ppm. It was also of interest to reduce
the amount of syngas H2 and CO lost to the amine solution. The conditions, mass flow rates,
and compositions of the entering and exiting streams are presented in Table 30. Table 31 shows
the percent recoveries of each of the components in the feed gas. Table 32 includes important
specifications of the amine scrubbing system regarding the absorber. The components of
interest in this unit operation are MDEA, H2O, H2, CO, CO2, H2S, NH3, and HCl. These tables
indicate that most of the H2S was absorbed out of the feed gas. The composition of the sweet
gas is less than 1 ppm H2S at 0.6799 ppm, which met specifications. 73.82% of the CO2 was
absorbed into the lean amine. In addition, over 99.99% of the H2 and CO was kept in the sweet
gas. These results indicated that the MDEA solution had little affinity for the syngas itself but a
high affinity for its contaminants. In addition, little MDEA (<0.0002%) was taken into the gas
stream; so, MDEA was not considered to leak out of the system in an appreciable amount to
the Aspen PLUS simulation. Some water was pulled into the sweet gas – this additional water
was considered in the Aspen PLUS simulation with the stream XTRA-H2O. Finally, it was
64
observed that all of the HCl was absorbed into the lean amine stream; this result made sense
because HCl is a strong acid that is more stable in the aqueous phase than the vapor phase. The
presence of this strong acid also solubilized over 99.999% of the NH3 into the aqueous phase.
Table 30: Stream specifications around the amine scrubbing absorber
Component flow rates (lb/hr)
Stream
T
P
m
MDEA
(°C)
(bar)
(lb/hr)
H2O
H2
CO
CO2
H2S
HCl
NH3
Feed Gas
20
35
5.52×104
0
49
6793
4.69×104
665.7
6.129
13.84
1.581
Sweet
28.5
35
5.47×104
0.010
102.4
6793
4.69×104
491.4
0.118
0
0
43.3
36
1.17×104
5275
6167
0
0
0
0
279.7
0
19.6
35
1.22×104
5275
6114
0.284
2.448
174.3
6.011
293.6
1.581
Gas
Lean
Amine
Rich
Amine
Table 31: Percent recovery (by mass) of each component in the feed gas stream
Mass recovery
H2O
H2
CO
CO2
H2S
HCl
NH3
126.180%
99.996
99.995
26.18
1.923
0
0.0001
(%)
Table 32: Important metrics for the amine scrubbing absorber
Sweet Gas H2S composition (ppm)
0.6799
Lean amine strength (wt% MDEA)
45
Lean amine temperature (°C)
43.33
Amine recirculation rate (barrel/day)
793.5
The rich amine was then throttled through a valve (Valve) to lower its pressure to 4.826 bar and
vaporize the volatile compounds in the rich amine. The valve was assumed to operate
adiabatically. The stream conditions for the inlet liquid (Valve) and outlet vapor-liquid mixture
(To Separator) are given in Table 33. This flashed mixture was then fed to an adiabatic
separator operating at 20.25°C and 2.413 bar. Light hydrocarbon gases (Light HC) leave the
65
separator at the top, and the remainder of the liquid (To Exchanger) leaves the bottom of the
vessel. These stream conditions are given in Table 33 below. This separator pulled 46% of the
H2 and 40% of the CO out of the liquid stream. The overall flowrate of the Light HC stream was
low because not many volatile hydrocarbons or gases existed in the aqueous phase.
Table 33: Stream specifications around the valve (Valve) and vapor-liquid separator (Separator)
Component flow rates (lb/hr)
Stream
T (°C)
P
m (lb/hr)
MDEA
H2O
H2
CO
CO2
H2S
HCl
NH3
(bar)
Rich Amine
19.6
35
1.22×104
5275
6114
0.284
2.448
174.3
6.011
293.6
1.581
To Separator
20.17
4.826
1.22×104
5275
6114
0.284
2.448
174.3
6.011
293.6
1.581
Light HC
20.25
2.413
1.32
0
0.015
0.130
0.970
0.136
0.002
0
0
To Exchanger
20.25
2.413
1.22×104
5275
6113
0.155
1.478
491.2
6.009
293.6
1.581
After passing through the separator, the rich amine stream is preheated in a shell-and-tube
heat exchanger (Lean/Rich Exchanger), where it is heated by the bottoms from the regenerator
(Regen Bottoms) to 93.33°C (Regen Feed). The stream specifications for the two streams
exchanging heat are noted below in Table 34 below. The hot stream (Regen Bottoms) provided
the cold stream (To Exchanger) with 1.406×106 Btu/hr of heat. According to Heuristics 26 and
31 presented by Seider, the minimum approach temperature was chosen to be 10°F (5.6°C) and
the pressure drop for each stream to be 5 psi (0.345 bar) (Seider, Seader, Lewin, & Widagdo,
2009). The temperature of Regen Feed was set to 93.33°C (200°F) because the reboiler duty in
the coming regenerator column (Regenerator) is generally at its smallest between 90 and 110°C
(Addington & Ness, n.d.).
66
Table 34: Stream specifications around Lean/Rich Exchanger
Component flow rates (lb/hr)
Stream
T (°C)
P
m (lb/hr)
MDEA
H2O
H2
CO
CO2
H2S
HCl
NH3
(bar)
To Exchanger
20.25
2.413
1.22×104
5275
6113
0.155
1.478
491.2
6.009
293.6
1.581
93.33
2.068
1.22×104
5275
6113
0.155
1.478
491.2
6.009
293.6
1.581
120.1
1.793
1.09×104
5275
5296
0
0
0
0
293.6
0
27.78
1.448
1.09×104
5275
5296
0
0
0
0
293.6
0
(cold in)
Regen Feed
(cold out)
Regen Bottoms
(hot in)
To Tank (hot
out)
The preheated stream Regen Feed was then fed to a regenerator distillation column
(Regenerator). The regenerator column provides heat to the rich amine solution, causing the
H2S, NH3, H2, CO, and CO2 to leave the column in the distillate. All the regenerated MDEA, all of
the HCl, and most of the water (86.6%) were recovered in the bottoms. A total reflux
condenser, composed of an air-cooled heat exchanger, a vapor-liquid separator, and a small
reflux pump, was used at the top of the column. The reboiler provides the duty to liberate the
acid gas species from the rich amine. The stream specifications around the column were found
and are reported in Table 35. An energy balance on the column was done with HYSYS, and the
resulting duties for the column are reported in Table 36. The column was assumed to operate
with no heat losses to the environment. The reflux drum (Regen Reflux Drum) and the reflux
pump (Regen Reflux Drum) depicted in the PFD were not simulated in the Aspen HYSYS
simulation, and so are not considered here.
67
Table 35: Stream specifications around Regenerator
Component flow rates (lb/hr)
Stream
T (°C)
P
m (lb/hr)
MDEA
H2O
H2
CO
CO2
H2S
HCl
NH3
(bar)
Regen Feed
93.33
2.068
1.22×104
5275
6113
0.155
1.478
491.2
6.009
293.6
1.581
Regen
120.1
1.793
1.09×104
5275
5296
0
0
0
0
293.6
0
105.0
1.517
1.32×103
0
817
0.155
1.478
491.2
6.009
0
1.581
Bottoms
Acid Gas
Table 36: Heat duties for Regenerator
Column Unit Operation
Energy Stream Name in HYSYS
Heat Duty (Btu/hr)
Condenser
QCondenser
-4.70×106
Reboiler
QReboiler
5.99×106
To provide the duty to the reboiler (QReboiler), a separate heat exchanger (Reboiler Steam)
was modeled in HYSYS. The purpose of this unit operation was to calculate the mass flow rate
of steam required to provide this heat. The steam was assumed to enter the reboiler as a
saturated vapor (Saturated Steam) and exit as a saturated liquid (Steam Condensate). The
steam rate, heat duty, temperature and pressure of the steam stream are reported in Table 37
below. It was noted that the heat duty of Reboiler Steam was equal to the duty of the reboiler
in the regenerator column (QReboiler) – this observation verifies that this steam rate provides
the necessary heat.
Table 37: Conditions of the saturated steam fed to Reboiler Steam
Steam rate (lb/hr)
T (°C)
P (bar)
Heat duty (Btu/hr)
6.56×104
147.7
4.461
5.99×106
The bottoms from the regenerator column (Regen Bottoms) then flowed through the Lean/Rich
Exchanger as described in the previous section. The cooled lean amine stream (To Tank)
entered a makeup block calculator (Surge Tank), where the additional water (Water Makeup)
and MDEA (MDEA Makeup) required to obtain an amine strength of 45 wt% was calculated. The
water and MDEA were assumed to be fed at the temperature and pressure of the To Tank
stream: 27.8°C (82°F) and 1.45 bar (21.0 psia), respectively. A purge stream (Makeup Purge)
68
was included with the makeup block but did not yield any flowrate. Another purge stream
further downstream was utilized instead. The flow rates, compositions, and temperatures for
the makeup block are given in Table 38: Stream specifications around Surge Tank below. Upon
mixing the additional MDEA and water, the temperature of the stream increased because of
the neutralization reaction of MDEA and HCl and the enthalpy of mixing of the water and
MDEA, as calculated by the Acid Gas Cleaning package in HYSYS.
Table 38: Stream specifications around Surge Tank
Component flow rates (lb/hr)
Stream
T (°C)
P
m (lb/hr)
MDEA
H2O
H2
CO
CO2
H2S
HCl
NH3
(bar)
To Tank
27.8
1.45
1.09×104
5275
5296
0
0
0
0
293.6
0
Water Makeup
27.8
1.45
1.20×103
0
1196
0
0
0
0
0
0
MDEA Makeup
27.8
1.45
2.77×102
277
0
0
0
0
0
0
0
Makeup Purge
-
-
0
-
-
-
-
-
-
-
-
To Pump
54.8
1.45
1.23×104
5552
6492
0
0
0
0
293.6
0
The newly recharged lean amine solution (To Pump) was then fed to a pump (Booster Pump) to
pressurize the stream to 36.34 bar. The temperature of the stream increased slightly from
54.8°C to 55.4°C as a result of the enthalpy change imparted by the pump:
=∆
= .
×
/ℎ (WPump in HYSYS). This stream (To Purge) was then sent to a splitter (HCl Purge),
where 5% of the flow (Purge) was diverted out of the simulation. This purge was implemented
to allow for convergence of the HYSYS simulation. Without this purge, HCl would have had no
outlet from the system and would have accumulated to very high levels, inactivating the amine.
The other 95% of To Purge (To Cooler) stayed in the regeneration cycle. Results of a material
balance around this splitter are shown in Table 39 below.
Table 39: Material balance around HCl Purge
Stream
Mass flow rate (lb/hr)
To Purge
1.23×104
HCl Purge
6.16×102
To Cooler
1.17×104
69
The stream To Cooler was then fed to a cooler (Cooler), where it flowed countercurrently to
chilled water in a heat exchanger. This unit operation functioned to lower the temperature of
the lean amine solution from 55.4°C to 44.3°C in order to prepare it for feeding to the absorber.
The pressure drop for the process fluid was chosen to be 5 psi (0.345 bar) (Seider, Seader,
Lewin, & Widagdo, 2009). A recycle block in HYSYS was used to match the properties of the
streams To Recycle and Lean Amine. The duty provided by the cooler (QCooler in HYSYS) was
found to be
= .
×
/ℎ . The lean, regenerated amine was then fed back into the
absorber, and the regeneration started again. The temperatures, pressures, compositions, and
mass flow rates of each stream around the cooler and recycle block are given in Table 40 below.
Table 40: Stream specifications around Cooler and RCY-1 (recycle block)
Component flow rates (lb/hr)
Stream
T
P
m (lb/hr)
MDEA
H2O
(°C)
(bar)
To Cooler
55.4
To Recycle
Chilled water
H2
CO
CO2
H2S
HCl
NH3
36.34
1.17×104
5275
6167
0
0
0
0
279.7
0
44.3
36
1.17×104
5275
6167
0
0
0
0
279.7
0
7.22
1.013
4.32×103
0
4320
0
0
0
0
0
0
32.2
1.013
4.32×103
0
4320
0
0
0
0
0
0
(in)
Chilled water
(out)
Material Balance
To check the closure of the system, an overall atom balance was performed on the Aspen
HYSYS simulation. The duties of the unit operations that cross the system boundary in the
simulation were also summarized. The results of this atom balance are Table 41 below. A
summary of the duties entering and leaving the system in this simulation is presented in Table
42 below. The discrepancies in the difference column in Table 41 were attributed to the recycle
block in the HYSYS simulation. These numbers could be reduced further by decreasing the
tolerance of the solver present in the recycle block.
70
Table 41: Overall atom balance for the amine scrubbing HYSYS simulation
Element
IN (lbmol/hr)
OUT (lbmol/hr)
DIFFERENCE (lbmol/hr)
H
7682.898
7682.903
-0.005
O
2143.672
2143.652
0.020
C
1714.118
1714.131
-0.013
S
0.180
0.180
0.000
Cl
0.380
0.403
-0.023
N
39.133
39.136
-0.003
Heat Duty
Table 42 provides a summary of the heat duty for the amine scrubbing system.
Table 42: Summary of duties which enter and leave the Amine Scrubbing HYSYS simulation (negative = output, positive = input)
Unit Operation
Regenerator (Reboiler)
Regenerator (Condenser)
Cooler
Booster Pump
Duty (Btu/hr)
6.00×106
-4.70×106
-1.94×105
2.36×104
Product Separation & Post-Processing
Process Description and PFD
Shown below in Figure 10 is the process flow diagram for dirty methanol separation and
purification. This is a key step in the final stages of the process to achieve the required
methanol purity of 99.97% in the product stream. A simulation was created in Aspen HYSYS
using the Peng-Robinson fluid package to model this process. A separator is used to produce a
liquid stream with few impurities that is fed to a distillation column, where the difference in
volatilities between the light key (methanol) and heavy key (water) is leveraged to obtain a
sharp separation.
71
Figure 10: Methanol purification process flow diagram
The process begins with the dirty methanol stream (RAW-METH) from the VL-SEP unit
operation entering an expansion valve (VLV-100) at 50°C and 35bar and exiting with the same
molar flowrate at 50.6°C and 1 bar. The valve lowers the pressure to allow for maximum gasliquid separation in the next unit operation; V-100 separator. 99.7% of the inlet stream (2PSFEED) leaves as liquid in stream (DISTILL-FEED) to be fed to the distillation column, while 0.3%
of the gas stream (WASTE-VAP) exits the separator at the top of the unit. The separation in V100 creates a much purer stream entering the distillation column; all of the gaseous impurities
such as trace amounts of carbon monoxide, carbon dioxide, ethylene, methane, and hydrogen
leave in the WASTE-VAP stream.
DISTILL-FEED enters the distillation column at 50.59°C and 1 bar with a composition of ~80%
light key (LK) and ~20% heavy key (HK). The LK component, methanol, has a boiling point of
66°C at 1 bar, while the HK component, water, has a boiling point of 100°C at 1 bar (Engineering
Toolbox, n.d.). Due to the relative volatility differences between the HK and LK, as shown in the
difference in boiling points, a sharp separation can be achieved. A full distillation column was
simulated in Aspen HYSYS with a total condenser operating at 2bar and a partial reboiler
operating at 2.758bar. The distillate vapor stream to the condenser is comprised of 99.97%
72
methanol and 0.02% water. Chilled water enters the condenser at 7.22°C and 1 bar to condense
the distillate vapor. The reflux drum splits the stream and a fraction of the condensate is sent
back to the column while another product stream (MeOH) leaves as pure liquid methanol. The
exit stream accounts for 58.3 million U.S gallons of methanol annually.
The liquid in the last stage of the distillation column is comprised mainly of the heavy key
component, 95.24%, with a small amount of dissolved light key, 4.76%. To recover more of the
light key component, this stream is sent to a reboiler. Low-pressure steam enters the reboiler
and easily vaporizes the methanol, which is sent back to the column. Consequently, the exit
stream increases its purity and additional methanol is injected back to the process to be
purified. The exiting stream (BOTTOMS) has 97.26% water and 2.74% methanol.
Table 43 lists a summary of operating temperatures, pressures, and mass flow rates entering
and exiting the column.
Table 43: Operating conditions and stream information for distillation column
Component flow rates (lb/hr)
Stream
DISTILL-FEED
MeOH
Bottoms
T (°C) P (bar) m (lb/hr)
50.59
80.04
1.0
2
5.07×104
4
4.471×10
CH4
H2O
H2
CO
0.6602 5.75x103 0.0741 1.7706
0.6602
4.4093
0
5.75x103
129.8 2.758 6.03×103
0.0741 1.7705
0
0
CO2
CH3OH
N2
0.4219 4.5x104 0.6415
4
0.4219 4.47×10 0.6415
0
287.61
C2H4
0.0055
0.0055
0
0
Material Balances
To check the closure of the system, an overall material balance was performed using the mass
flow rates in lbmol/hr given in the Aspen HYSYS simulation. To do this, an atomic species
balance was calculated where the inlet stream, Raw-MeOH, was compared to the outlet
streams MeOH, Bottoms, and Waste-Vap. Table 44 displays the material balance calculations
around the entire post-processing system.
Table 44: Overall atom balance for MeOH Purification system in HYSYS
Element
IN (lbmol/hr)
OUT (lbmol/hr)
DIFFERENCE (lbmol/hr)
H
6268.279
6268.608
0.329
O
1727.299
1727.391
0.092
73
C
1408.131
1408.298
0.167
N
0.482
0.482
0
TOTAL
9404.19
9404.778
0.588
Heat Duty
The entering and exiting heat duties were evaluated for the methanol post-processing section.
In this section, all the required duties are supplied via chilled water or low-pressure steam. An
in-depth analysis of the utility cost is included in section IX. Shown below in Table 45 is the
summary of the energy balance per unit operation, condenser and reboiler, where positive
duties are considered inputs to the system and negative duties are outputs from the system.
Table 45: Summary of duties which enter and leave the MeOH Purification HYSYS simulation (negative = output, positive = input)
Unit Operation
Reboiler
Condenser
Duty (Btu/hr)
4.088×107
-3.743×107
74
VIII. Process Description & Equipment Specifications
Equipment design for the solar-thermal biomass gasification facility was accomplished by
dividing the facility into subgroups based on function and end-product. These sections include
biomass pre-processing, methanol gasification, methanol purification, amine scrubbing and
solar field design. Biomass pre-processing takes in corn stover as feed, which is ground and
separated into usable biomass suitable for gasification. The primary plant section is the
methanol gasification system, in which this usable biomass is reacted to form raw methanol.
The raw methanol stream then enters the methanol purification system, in which a series of
separators remove impurities, and result in a high-purity methanol product stream. The amine
scrubbing system removes toxins and environmental hazards from the gas stream exiting the
methanol gasification subsystem. The solar field provides power to solar reactor, which
provides a means for driving a portion of the methanol gasification process.
Equipment design specification begins with a section dedicated to generalized equipment
design, in which the procedure for designing common pieces of process equipment that are
present throughout the facility is elucidated in detail. Comprehensive equipment design
calculations may be found in Appendix II-A through Appendix II-L.
Generalized Equipment Design
Pumps
Pumps are present in both the methanol gasification section and in the amine scrubbing
system. In these processes, pumps function to raise the pressure of a process stream prior to
that stream entering a reactor, separator, or other piece of process equipment that requires
input at a specified pressure.
For all pumps, specifications were found in concordance with the Carver RS Series Technical
Support Information pump design manual (Carver, 2006). This technical support manual
provides multiple pump sizing options, based on flowrate (gpm) and total feet of head in each
case.
The volumetric flowrate of each process stream ( ̇ ) was obtained directly from Aspen Plus or
HYSYS simulations, while the total feet of head had to be calculated. Required total pressure
75
head (�
) was determined by dividing the desired pressure difference in the process stream,
∆ , by the stream density, . Here, pressure units of pounds per square foot
⁄
density units of pounds per cubic foot
desired by the product manual.
�
⁄
and
returned pressure head units in feet, as
= ∆ ⁄
(30)
The pressure head requirements and volumetric flowrates for each stream were then used in
conjunction with the Carver Pump Company pump curves. An example pump curve from the RS
Series manual is provided below in Figure 11.
Figure 11. Carver Pump Company pump curve for Size A low flow, high RPM centrifugal pump (Carver, 2006).
These pump curves were used by finding the point on the curve at which the total head and
flowrate for a process stream intersect. This point corresponded to an impeller size, listed in
inches above the solid black lines, as well as a horsepower requirement, listed to the right of
the straight dashed lines. Pump efficiency was determined by following the grey semicircular
76
lines with percent efficiencies indicated. The pump size, RPM value and casing size are provided
in the black box in the top right corner of each pump curve.
Often, the total pressure head required for a given process stream was several factors higher
than the total head specified in the manufacturer-provided pump curves. In this case, the total
pressure head (�
) calculated by Equation (30) was divided by an integer number that yields
a head value present on the provided pump curves (�
the number of stages required for that pump (
�
=
�
). This integer number represented
).
⁄
(31)
Final considerations for pump design included ensuring that the available net positive suction
head, NPSHA, exceeded the required net positive suction head, NPSHR, for each design. For
pumps involved in the HYSYS simulations, NPSHA was obtained directly from the software. For
Apsen PLUS simulations, NPSHA must be hand-calculated by Equation (32):
=
−ℎ −ℎ −
(32)
In this equation, ℎ and ℎ refer to the pump elevation and pump losses respectively. Because
no pumps in the process had specified elevations, because head losses were assumed to be
minor (and therefore neglected in NPSHA calculations), both the elevation and loss terms
reduce to zero. The available net positive suction head equation then reduced to the
atmospheric pressure and vapor pressure divided by their respective specific weights. This
simplification is shown in Equation (33):
=
−
(33)
Due to the low vapor pressures and high inlet pressures of the process streams involved in each
simulation, the available net positive suction head for each pump greatly exceeded the NPSHR
indicated by the equipment manufacturer.
77
Material and Energy Balances
The pumps specified in this process were assumed to operate at steady state, in which the mass
entering is constant and equal to the mass leaving each piece of equipment. Because of this,
the mass balance for each pump was trivial.
The energy balance associated with each pump can be expressed by the following equation:
∆
+
−
�
=
∙
(34)
̇
In Equation (34), the pressure head, ∆ ⁄ is added to the average squared velocity of the inlet
and outlet streams. This term was then set equal to the shaft work,
, multiplied by the pump
efficiency, , and divided by the mass flow rate passing through the device, ̇ . This expression
is a simplified version of the centrifugal pump shaft work expression found in Fundamentals of
Fluid Mechanics, 7th Ed. (Munson, Heubsch, & Rothmayer, 2012). Because the inlet and outlet
streams experienced negligible vertical height change, the term accounting for fluid rise is
omitted.
Shell-and-Tube Heat Exchangers
The majority of heat exchangers present throughout the biomass gasification facility were
modeled as shell-and-tube heat exchangers. Shell and tube heat exchangers consist of a shell
side and tube side. A fluid passes through either side, and heat transfer is encouraged between
them. Usually, the more corrosive fluid passes through the tubes, while the more inert fluid
passes through the shell. This choice is made for both practical and economic purposes; it is
easier to manually clean tube sheets than it is to clean the shell casing of a heat exchanger.
Furthermore, if a more durable, expensive material is required for a corrosive fluid, it is less
costly to make only the tubes from this high-cost material than to make both the tubes and
shell casing from a high-end substance (Seider, Seader, Lewin, & Widagdo, 2009). A schematic
of a simple shell-and-tube heat exchanger is shown in Figure 12 below.
78
Figure 12. Schematic of a shell-and-tube heat exchanger.
In order to specify heat exchanger parameters, the heat duty required for the exchanger must
first be determined. This value was obtained directly from HYSYS or Aspen PLUS and was
recorded in units of Btu/hr.
The hot and cold streams in the heat exchanger were then identified, and an appropriate utility
was selected. Heating utilities included, but were not limited to, high pressure or low pressure
steam; cooling utilities included cooling water or refrigerated brine (Seider, Seader, Lewin, &
Widagdo, 2009). Occasionally, the designed heat exchangers used two process streams rather
than a process stream and a utility stream. When this arrangement was possible, heat transfer
without the need for additional utility costs was achieved.
The mass amount of utility required to achieve the desired heating or cooling of the process
stream was either be obtained directly from the process simulation, or calculated by hand using
an energy balance, such as the one shown in Equation (35) below:
where
=
̇
(35)
∆
is the specific heat capacity of the utility and ∆ is the difference between the inlet
and outlet temperatures of the utility. If the utility stream involved condensation (e.g. of
steam), ∆
, the heat of vaporization of the utility, was used instead of
∆ .
79
The primary parameter used in characterizing shell and tube head exchangers is the heat
exchanger area (
). This area is calculated by the following expression:
=
∆
̇
(36)
In Equation (36), ̇ refers to the process enthalpy in units of Btu/hr.
is a heat transfer
coefficient unique to each pair of fluids undergoing heat transfer. These values can be found in
standard tables available in Pe y’s Che i al E gi ee i g Ha d ook (Green & Perry, 2008) and
Product and Process Design Principles (Seider, Seader, Lewin, & Widagdo, 2009), and have units
of Btu/(hr·ft2·°F). The term ∆
is the log-mean temperature difference, defined by the
following equation:
∆
=
(
ln
ℎ
(
,�
ℎ
,�
−
,
−
,
)−
)
⁄
ℎ
,
ℎ
,
−
−
,�
(37)
,�
After substituting Equation (37) into Equation (36), the final result is a heat exchanger area with
units of ft2.
In order to determine the precise heat exchanger properties required in each case, initial
estimates were required for specific parameters. Using heuristics obtained from Product and
Process Design Principles, initial estimates for tube outer diameter and tube length were set at
0.75 in. and 16 ft, respectively (Seider, Seader, Lewin, & Widagdo, 2009). The flow area per
tube was determined by cross-referencing the tube O.D. with a selected tube wall thickness.
The typical thickness was 14 BWG; this was a heuristic estimate (Seider, Seader, Lewin, &
Widagdo, 2009).
The area per tube (
) was then calculated by the following formula:
=
∙
. .
∙
(38)
In Equation (38), the tube O.D. was first converted to units of ft.
The number of tubes per heat exchanger was then determined by dividing the total heat
exchanger area, , by the area per tube:
80
= /
(39)
In order to determine if the initial heuristics and estimates are valid for a given case, the tubeside velocity, , is calculated by the following equation:
= ̇ ∙( )∙(
)∙( )
(40)
A desirable tube-side velocity is between 1 and 10 ft/s. In order to ensure that flow through the
heat exchanger tubes met this criterion, the tube O.D., tube length and BWG were adjusted
until this parameter was met. Finally, Table 18.6 in Product and Process Design Principles was
used to identify tube spacing, number of shell and tube passes, and shell internal diameter.
Material and Energy Balance
Because all heat exchangers involved in this process were assumed to operate at steady-state,
the mass that entered both the tube and shell sides of each heat exchanger was equal to the
mass that leaves. For this reason, the material balances for each heat exchanger were trivial,
but were important in achieving the desirable tube-side velocity.
The design of each heat exchanger was ased a ou d the heat e ha ge s e e g
ala e. This
energy balance was given in the following equation:
,
−
,�
=
ℎ
,�
−
ℎ
,
(41)
In Equation (41), the enthalpy decrease in the hot fluid is equal to the enthalpy increase in the
cool fluid. In this way, all heat exchangers were assumed to be perfectly insulated.
81
Vapor-Liquid Separators
In this process, there were numerous instances where vapor-liquid mixtures were separated
into their respective vapor and liquid constituents. Traditional gravity-settling separators were
used in the entirety of this process design because of their simple construction and robustness
of operation (Stewart & Arnold, 2008). Horizontal and vertical varieties of these separators are
composed of four primary sections; schematics of these pieces of equipment are given in Figure
13 and Figure 14, respectively. The first section is an inlet diverter, which suddenly changes the
di e tio of i let flo a d e a les [a ] i itial g oss sepa atio of the li uid a d gas phases
(Stewart & Arnold, 2008). The second section is a gravity settling section, which provides a
space for small liquid droplets to settle out of the vapor phase. In this process, all separators
were sized to separate out liquid droplets of diameters between 150 and 200 µm. Larger liquid
droplets outside of this range can clog the third section, the mist extractor. This component is a
esh that p o ides a la ge su fa e a ea to
oales e a d e o e the s all d oplets of li uid
that remain suspended in the vapor phase (Stewart & Arnold, 2008). The fourth section is a
liquid collection section, which serves to drain out the liquid phase from the vessel. The gas
outlet is equipped with a pressure control module and pressure control valve to regulate
pressure and monitor constriction of the mist extractor. The liquid outlet is fitted with a level
meter and level control valve to ensure the liquid level remains at the design liquid level.
82
Figure 13: Schematic of a horizontal gravity-settling vapor-liquid separator (Stewart & Arnold, 2008)
Figure 14: Schematic of a vertical gravity-settling vapor-liquid separator (Stewart & Arnold, 2008)
83
The design procedure for gas-liquid separators outlined below comes from chapter 3 of the
book Gas-Liquid and Liquid-Liquid Separators by Stewart and Arnold (Stewart & Arnold, 2008).
The procedures for the horizontal and vertical units vary slightly, but some information about
the liquid droplets must be determined regardless of the alignment of the unit. All
computations in these designs were done with variables in field units. The variables used in this
section are summarized in the table of nomenclature (Table 46) below. All separators were
sized to be half full of liquid to match the theory presented in the aforementioned text.
Table 46: Table of nomenclature for gas-liquid separator design
Parameter
Symbol
Gas compressibility
Gas flow rate
Liquid flow rate
Operating temperature
Operating pressure
Drag coefficient
Liquid droplet diameter
Units
Comments
-
Evaluated at outlet conditions
MMSCFD
Evaluated at outlet conditions
bbl/day
Evaluated at outlet conditions
°R
Evaluated at outlet conditions
psia
Evaluated at outlet conditions
-
-
µm
-
Mass density of gas
lb/ft3
Evaluated at outlet conditions
Mass density of liquid
lb/ft3
Evaluated at outlet conditions
Vessel internal diameter
in.
-
cP
Evaluated at outlet conditions
ft/s
-
min
-
-
-
ft
For horizontal units
ft
For horizontal units
ft
Larger value of
Slenderness ratio
-
-
Seam-to-seam length
ft
-
in.
For vertical separators
Viscosity of the gas
Liquid droplet settling velocity
�
Retention time of the liquid
Reynolds number
Gas capacity length
,
Liquid capacity length
,
Overall effective length
Height of liquid volume
ℎ
,
,
,
84
The settling velocity
drag coefficient
of the liquid droplets to be separated was first determined to find the
of the droplets. To do so, a recursive approach was used. First, a guess for
was found using Equation (42):
The Reynolds number
(
[
= .
−
)
]
/
(42)
for the gas flow around the droplet is then found with Equation (43):
= .
(43)
�
From here, the first approximation of the drag coefficient
was calculated using a correlation
that applies for turbulent and laminar flow around the droplet. This correlation is shown in
Equation (44):
A better approximation for
=
+
/
+ .
−
)
(44)
was then computed using Equation (45):
(
[
= .
]
/
(45)
From here, Equations (43)-(45) were iterated until convergence was achieved in Excel. An
estimate for the liquid retention time
was determined from Table 47 below, provided in the
aforementioned Stewart and Arnold text.
Table 47: Retention time for two-phase separators – Table 3.2 in (Stewart & Arnold, 2008)
85
Horizontal Separators
Horizontal separators are commonly used in the field for two reasons. First, they are often
s alle a d thus less e pe si e tha a e ti al sepa ato fo a gi e gas a d li uid flo
ate
(Stewart & Arnold, 2008). This suggestion makes sense, as vertical separators are more
sensitive to wind and earthquakes, requiring vertical separators to be of thicker construction
than their horizontal counterparts. The liquid-gas interface is also larger in a horizontal vessel,
allowing more surface area for bubbles of gas to rise into the gravity settling section and
droplets of liquid to fall into the liquid collection section. Second, horizontal separators are
o
o l used i flo st ea s ith high gas–li uid atios (Stewart & Arnold, 2008). This
heuristic makes sense, as not much liquid volume and thus vessel volume is required.
The gas capacity constraint was considered to determine the effective length
required
,
for the liquid droplets to separate from the gas phase. This value was determined using
Equation (46):
,
=
[
][
(
−
)
]
/
(46)
In addition, a liquid capacity constraint was investigated to provide an effective length
,
needed so that the liquid and gas can reach vapor-liquid equilibrium. This value was
determined using Equation (47):
,
=
.
Notice that these expressions are both a function of the vessel internal diameter
(47)
and the
effective lengths for gas and liquid. Therefore, to determine each of these values, a set of
standard diameters
and
,
in increments of 6 in. were assembled in a table, and the values
,
calculated for each . For each , the larger effective length value is taken as the
design effective length
. To make a final decision on which combination of
used, a slenderness ratio
was computed according to Equation (48):
=
and
was
(48)
86
Following the recommendations of the text, a design was chosen such that
<
< . This
decision was made in the interest of reducing cost and ensuring no waves form in the vessel, reentraining droplets in the vapor (Stewart & Arnold, 2008). Finally, the effective length was
translated into an actual seam-to-seam length
to provide space in the vessel for the mist
extractor, liquid outlet, and inlet diverter. A schematic of this length is provided in Figure 15
below. As a final consideration, the final design was chosen such that the seam-to-seam length
did not exceed the relation given in Equation (49):
=
(49)
Figure 15: Approximate seam-to-seam length for a horizontal separator operating half-full (Stewart & Arnold, 2008)
These design calculations provided the metrics for the internal diameter and seam-to-seam
length of the vessel. Other metrics, especially material of construction and thickness of the
outer shell, were determined through the design of pressure vessels in the following section of
this report. The aforementioned calculations are unique for gas-liquid separators.
Vertical Separators
Ve ti al sepa ato s a e
o
o l used i flo st ea s ith lo to i te
ediate gas–liquid
atios (Stewart & Arnold, 2008). The liquid-gas interface is smaller than in a horizontal vessel,
87
and this quality is acceptable because less surface area is needed for bubbles of vapor to rise
into the gravity settling section and droplets of liquid to fall into the liquid collection section.
The design procedure for the vertical separators was similar to that for the horizontal
separators. First, the gas capacity constraint was developed. In this setup of gas-liquid
is the gas apa it
separator, the vessel diameter
allow liquid droplets to separate f o
Unlike with the horizontal separator,
o st ai t, a d it must be maintained to
the e ti all
o i g gas (Stewart & Arnold, 2008).
was determined with a function (Equation (50) below)
that is not also a function of separator length:
=
[
][
(
−
)
]
/
(50)
Once the diameter was found, a liquid capacity constraint was calculated. Here, the liquid
capacity constraint is the height of the separator ℎ and must be large enough to provide the
liquid retention time required for the vapor and liquid to reach equilibrium. ℎ was computed
using Equation (51) below:
Finally, the seam-to-seam length
ℎ=
.
(51)
was determined. For a vertical separator, enough space
must be allocated for the mist extractor (about 6 in.), the gravity settling section ( + 6 in. or 42
in. minimum) and a small clearance for the liquid outlet at the bottom of the unit. A schematic
of these specifications is given in Figure 16: Schematic for determining approximate seam-toseam length for a vertical separator below. There exists two equations in (Stewart & Arnold,
2008) to compute
– the larger value of those given by Equations (52) and (53) was used to
size the vertical separator.
=
=
ℎ+
ℎ+
(52)
+
(53)
88
Figure 16: Schematic for determining approximate seam-to-seam length for a vertical separator (Stewart & Arnold, 2008)
Finally, the slenderness ratio
was examined with Equation (48). A larger value for
reduces
the cost of the vessel because less construction material is used. However, a slenderness ratio
that is too high will result in a separator with an unreasonable liquid level height and a
substantial decrease in mechanical integrity for the same shell thickness. So, Steward and
Arnold recommend that the slenderness ratio be less than 4, which values between 3 and 4
being common (Stewart & Arnold, 2008). The diameter
implemented in the design was
chosen to satisfy this criterion.
89
Once again, these design calculations provided the metrics for the internal diameter and seamto-seam length of the vessel. The material of construction, thickness of the outer shell, and
number of trays (in the case of the absorber), were determined through the design of pressure
vessels in the following section of this report. The aforementioned calculations are unique for
gas-liquid separators.
Material and Energy Balances
In separator units, the material and energy balances are simple to express. The material
balance can be expressed according to Equation (54) as follows:
̇� = ̇
,
+ ̇
(54)
,
where, in the absence of chemical reaction, the inlet mass or molar flow rate ̇ � is equal to
the sum of the outlet gas flow rate ̇
,
and the outlet liquid flow rate ̇
stream may be either vapor, liquid, or a vapor-liquid mixture.
,
. The inlet
The energy balance for a separator unit can be more complex than its material balance. The
energy balance can be expressed according to Equation (55) as follows:
where
�
̇�
�
= ̇
,
,
+ ̇
,
,
+ ̇
represents the specific enthalpy (or molar enthalpy) of the inlet stream,
specific enthalpy (or molar enthalpy) of the outlet gas stream,
,
(55)
,
is the
is the specific enthalpy (or
molar enthalpy) of the outlet liquid stream, and ̇ is the heat duty of the separator. The heat
duty is nonzero if the separator must be cooled or heated to achieve the desired outlet
temperature.
Pressure Vessels and Towers
Pressure vessels and towers are a very common piece of equipment in chemical processing
pla ts, se i g a a iet of fu tio s i ludi g eflu d u s, flash d u s, k o k-out drums,
settle s, he i al ea to s,
i i g essels […] a d sto age d u s (Seider, Seader, Lewin, &
Widagdo, 2009). These vessels are usually mostly empty, with very little internals other than
skirts for support, nozzles, and manholes for internal access. Platforms and ladders are often
added to particularly large vessels for maintenance access. Pressure vessels can also be
90
oriented horizontally or vertically; the choice of orientation depends on the application. These
pieces of equipment can be made out of a variety of materials, the choice of which also
depends on application. An example of a horizontal pressure vessel can be seen in Figure 17.
Figure 17: An example of a horizontal pressure vessel (Pressure Vessels, 2014)
In this process design, cylindrical pressure vessels with hemispherical caps operating at positive
gauge pressures were designed. This design is common in many applications, so it was chosen
to be the standard in this report. In determining the specifications of a pressure vessel, a few
important pieces of information arising from the specific application of the vessel must be
known. These values, as well as all the variables involved in the vessel design, are given in the
table of nomenclature (Table 48) below.
Table 48: Table of nomenclature for pressure vessel design
Parameter
Internal diameter
Seam-to-seam length
Tangent-to-tangent length
Symbol
�
Units
Comments
ft
-
ft
-
ft
-
Operating pressure
psig
Assuming ambient pressure = 14.7 psi.
Design pressure
psig
Assuming ambient pressure = 14.7 psi.
Operating temperature
°F
-
Design temperature
°F
-
91
Maximum allowable stress
Density
Weld efficiency
psi
Of material of construction.
lb/ft3
Of material of construction.
-
-
Thickness to withstand internal pressure
in.
For all vessels.
Thickness to withstand wind and earthquakes
in.
Only for vertical vessels and towers.
Thickness without corrosion
in.
For all vessels.
Corrosion thickness clearance
in.
For all vessels as a safety measure.
Thickness of vessel shell
in.
For all vessels.
Outer diameter of vessel
ft
-
Weight of the vessel
lb
-
Before designing a pressure vessel or tower, a few variables from Table 48 were determined
fo
the u it s i di idual appli atio and the simulation of their operation in Aspen PLUS or
Aspen HYSYS: the internal diameter
the operating temperature
�,
the seam-to-seam length
, the operating pressure
,
, and the material of construction. Once this information was
obtained, the specifications could be calculated according to the method of Mulet, Corropio,
and Evans (Seider, Seader, Lewin, & Widagdo, 2009). First, the tangent-to-tangent length
was
estimated assuming the cylindrical vessel is equipped with 2:1 elliptical heads, as recommended
by Seider in Equation (56) (Seider, Seader, Lewin, & Widagdo, 2009):
=
+ .
�
If additional information about the tangent-to-tangent length
was used in lieu of Equation (56). Next, the design pressure
(56)
was known, that information
was calculated. The design
pressure is an estimation of the maximum pressure surge that might occur in the vessel during
normal operation. For operating pressures between 0 and 5 psig,
=
� . If the
operating pressure was above 1000 psig, the design pressure was chosen to be
Otherwise, the design pressure was determined using Equation (57):
= ex�{ .
The design temperature
+ .
[ln
]+ .
[ln
] }
= .
.
(57)
, an estimation of the maximum temperature that might occur in the
vessel during normal operation, was chosen to be 50°F higher than the operating temperature
. The maximum allowable stress
and density
were determined by the properties of the
92
material of construction at the design temperature. The weld efficiency
0.85 if the thickness of the shell
< 1.25 in.; otherwise,
to withstand the internal pressure
was chosen to be
= . The thickness of the vessel wall
was calculated from Equation (58):
�
=
(58)
− .
At lower design pressures, the value from Equation (58) was too small to achieve proper rigidity
of the vessel. If this was the case, a minimum wall thickness was chosen from Table 49 below:
Table 49: Minimum wall thicknesses of a pressure vessel to ensure sufficient rigidity (Seider, Seader, Lewin, & Widagdo, 2009)
Vessel inside
diameter (ft)
Up to 4
Minimum wall
thickness (in.)
1/4
4-6
5/16
6-8
3/8
8-10
7/16
10-12
1/2
If the pressure vessel was horizontal,
=
, the thickness of the vessel in the absence of
corrosion. If the pressure vessel was vertical or a tower, additional considerations were made
to correct for the effects of wind and earthquakes on the structural integrity of the unit.
Assuming a wind velocity of 140 mph acting on the entire vessel uniformly, Equation (59) was
used to determine the additional thickness required to withstand this stress
=
.
+
(59)
Notice that Equation (59) depends on the outer diameter of the vessel
determined yet. This value was estimated for a shell thickness
.
, which has not been
of 0.5 in. by geometry with
Equation (60) below, then reevaluated at the end of this procedure.
=
�
+
(60)
93
Next,
was found from the average of
, the thickness of the vessel at the top, and
the thickness at the bottom. A corrosion clearance
the final calculated
= ⁄ � . was then added to
+
,
to give
. Vessels are often fabricated in standard increments according to the
following assumed protocol from Seider (Seider, Seader, Lewin, & Widagdo, 2009):



1/16 increments for 3/16 to 1/2 in. inclusive
1/8 increments for 5/8 to 2 in. inclusive
1/4 increments for 2 1/4 to 3 in. inclusive
If the vessel needed to be thicker than 3 in., it was assumed that the vessel would have to be
custom-made with the calculated
. Otherwise, the value for
was rounded up to the
appropriate increment given above, and the procedure to calculate
(59) was repeated. This procedure was iterated until the final value of
the weight of the vessel
starting with Equation
was constant. Finally,
, an important parameter for costing and transportation
considerations, was determined with Equation (61) as follows:
=
�
+
(61)
This design procedure yielded two key specifications of the pressure vessels: their tangent-totangent length and the thickness of the shell. The weight of the vessel was also computed.
Cyclones
Centrifugal cyclones are commonly used in gas-solid separation processes with two main
functions: recover as much material as possible while minimizing contamination of material to
the outside environment. Many factors influence the type of cyclone employed and its
efficiency, such as particle size distribution, density of the particles, and inlet gas velocity.
Cyclones generally have a relatively simple operation, with no mechanical moving parts, and
little maintenance and thus low capital costs. Common types of cyclones include reverse flow,
straight-through flow, and impeller collectors (Flagan & Seinfeld, 1988). The reverse-flow
cyclone design was used in this design, as this is frequently utilized to remove fly ash and larger
particles. Particles smaller than 50µm are harder to separate and may require additional
equipment, such as a filter, which would also increase the cost. Although, with particle sizes
above 50 µm, efficiency upwards of 90% can be achieved with a basic cyclone unit (Flagan &
94
Seinfeld, 1988). Reverse or straight flow cyclones have a design advantage in that they can
accept higher gas velocities, and thus higher pressure drops. Their disadvantage is that solids
can build-up at the bottom collection of the unit. Cyclones can be made from a variety of
materials that are able to endure high temperatures and abrasive particles. Erosive wear and
tear can happen where materials can move easily within the unit from contact with particles
and surfaces, and where materials are loaded. Thus, extremely smooth surfaces are needed for
the interior of the cyclone which will minimize erosion and allow the particles to collect more
easily.
The cyclone operation relies on inertial deposition of solids from a gas stream. The gas-solid
stream enters the cyclone tangentially, where the stream initially flows downward in a spiral
manner. Due to the inertial force on the particles and the density difference between the
particles and the gas, the particles separate and collect on the outer surface of the cyclone and
fall to the bottom collection. The clean gas stream then reverses direction and flows upward
out of the cyclone unit. Figure 18 displays the basic operational scheme in a reverse flow
cyclone.
Figure 18: Basic operation of a reverse flow cyclone
95
In this plant design, high-efficiency cyclones were chosen based on the desire to achieve
maximum separation between the gas and solid particles. The following explanation details the
design of cyclones. The entering stream is a mixture of gaseous and solid phases at 121.8°C and
35bar with a volumetric flowrate of
ℎ
, where the vapor fraction by mass is 0.955 and
the solid fraction by mass is 0.002. The solids to be separated are solid ash from the solar
reactor, and the gases are comprised of a variety of hydrocarbons, nitrogen, oxygen, water
vapor, chlorine and sulfur. The particle size distribution is estimated to be between 100-200µm.
The density of the gas is approximately .
.
, while the density of the particles is
. Figure 19 shows a high efficiency cyclone with standard dimensions.
Figure 19: Dimensions of a high-efficiency cyclone
The area of the inlet duct tube was first calculated using Equation (62) below, where
volumetric flowrate of the inlet stream and
is the
is the optimum inlet velocity given to be 15 m/s
(James R. Couper, 2012):
=
(62)
96
The cyclone diameter,
, can be calculated from the inlet duct area, which is shown below in
Equation (63).
=√
(63)
.
A standard cyclone diameter is 0.667 feet. The calculation from Equation (63) yielded a design
diameter of 2.12 feet, which requires splitting the design into 4 cyclones in parallel, thus
producing a diameter of 1.05 feet. A scaling factor is then determined using Equation (64),
where the variable nomenclature is given below in Table 50:
= [(
) ∗(
)∗
∆�
∆�
∗(
�
)]
�
(64)
Table 50: Table of nomenclature for cyclone design equations
Parameter
Symbol
Units
Value
Standard cyclone diameter
in
8
Design cyclone diameter
in
calculated
Standard volumetric flowrate
ft3/hr
7869.12
Design volumetric flowrate
ft3/hr
calculated
∆�
lb/ft3
124.85
∆�
lb/ft3
calculated
�
cP
0.018
�
cP
calculated
Solid-fluid density difference at standard conditions
Design solid-fluid density difference
Test fluid viscosity at standard conditions (air at 1atm, 20°C)
Design fluid viscosity (H2 at 35 bar, 121.8°C)
Mean diameter of particle separated at standard conditions
N/A
Design mean diameter of particle separated
Design scaling factor
N/A
⁄
N/A
calculated
By calculating a scaling factor of 2.38, the final cyclone height and diameter was determined to
be 11.3 feet with a diameter of 2.5 feet. The material chosen was Monel, as the gas stream
constituents contain corrosive compounds. If the i let st ea
does t o tai a
o osi e
materials, then a more affordable material can be chosen that will withstand the temperature
in the cyclone, such as carbon steel.
97
Material Balance
A material balance was performed around the cyclone to ensure mass was conserved in the
separation process. Molar flow rates were given from the Aspen PLUS simulation. Aspen
assumed the cyclone to operate at 100% efficiency, thus all of the input molar flows for each
compound equaled the output molar flow for that compound. The cyclone does not operate
with a heat duty, therefore an energy balance is not applicable.
Biomass Pre-Processing
P-1 / SR-101 Shredding
The first unit operation in the biomass pre-processing section is the stage 1 size reduction. This
is accomplished by a shredding machine. The required outlet particle size was to be no larger
than ′′ diameter. The feed into this process will be corn stover in bale form with a mass
flowrate of .
ℎ
with a water content of 25.0%. The bales will enter at 25°C and 1.013
bar. Important design specifications for size reduction equipment are the desired particle size,
power requirement, capacity, as well as operating conditions such as pressure and
temperature. Listed below in Table 51 are the requirements, which were determined from the
simulation carried out in SuperPro Designer.
Table 51: Required Design Specification for P-1 / SR-101 Shredding
Design Specification
Value
¼ �
Maximum Particle Size
Power Requirement
Capacity
Pressure
Temperature
.
.
.
ℎ
ℎ
ℎ
°
Using these design specifications, multiple options were considered, but a rotary knife cutter
was determined to best fit the required design specifications. This unit is designed by S. Howes,
Inc. based out of Silver Creek, NY (S. Howes, Inc., 2013). The rotary knife cutter specifications
are shown below in Table 52. A rotary knife cutter is a large rotary drum with fixed and rotary
98
bed knives powered by a belt or clutch drive. A sizing screen is used to prevent oversized
material from exiting before desired size reduction has been achieved.
Table 52: Design Specifications for Model KC11 Rotary Knife Cutter
Model
Capacity (lbs/hr)
KC11
−
RPM
Power (Btu/hr)
.
Weight (lbs)
,
Using these design specifications, it is noted that one unit does not have the requisite capacity
for the desired mass flowrate, so a total of 4 units must be used to accomplish this. They will be
run in parallel. The total power requirement using these options is slightly under specification,
with 4 units providing a total power of .
x
B
hr
. The manufacturer has the option of
installing a clutch drive transmission instead of the standard belt driven option and this is
assumed to be adequate for matching the power requirements such that the machines will be
able to meet the power demand. The screen size necessary to achieve the maximum particle
size above is Standard Mesh 3.5, which has an opening of 0.25 inches. The knife cutter layout is
shown below in Figure 11 and the dimensions are specified in
Table 53.
Figure 20 : Rotary Knife Cutter Layout with Dimensions
99
Table 53: Length Specifications for Rotary Knife Cutter to be used with Rotary Knife Cutter Layout
Model
Length A
Width B
Number
(in)
(in)
KC11
36-1/2
26
Base D x E in
Inlet F x G in
37-5/8 x 60
5-1/2 x 22
The feed as well as the outlet from the rotary knife cutters will be on belt conveyers as they are
economically efficient and the particles have a large enough diameter such that enclosure is not
necessary at this point of the process. The materials of construction are not specified on the
specification sheet provided by the manufacturer, yet at this point, there are no requirements
for specialty materials or design considerations so it is assumed that the material used will be
adequate for this unit operation.
P-2 / RDR-101 Rotary Drying
The next step of the biomass pre-processing section is to reduce the moisture from 25% to
6.25%. This is modeled as a rotary dryer, which is a large cylinder that is horizontal to the
g ou d ith a s all a gle. It otates alo g its e t al a is ith affles, o flights, that help to
mix and move the solid material. To accomplish the drying, a counter-current hot air stream
entering at 120°C and exiting at 90°C is used. The duty of the dryer is found in the simulation to
be .
/ℎ and the biomass exits at 90°C with a flowrate of .
rotary dryer is shown below in Figure 21 (M. G. Silva, 2012).
ℎ
. A typical
Figure 21: Conventional Rotary Cascade Dryer
100
Important design considerations for rotary dryers include length, diameter, residence time,
number of flights, and revolutions per minute. To design the rotary dryer for this process,
example 9.6 is used in Chemical Process Equipment: Selection and Design (James R. Couper,
2012). The table of nomenclature is shown below in Table 54.
Table 54: Table of Nomenclature for Rotary Dryer
Symbol
Description
Amount of dry air per unit area
allowed
Diameter of dryer
Mass flowrate of air
̇
Volume of dryer
Length of dryer
Number of flights
RPM of dryer
�
Residence time
Cross-sectional area
Bulk density, corn stover
Area % occupied by solids
Unit
ℎ
ℎ
�
�
ℎ
�
%
From the simulation in SuperPro Designer, the air flowrate necessary, ̇ , was .
ℎ
. To
determine the size of the dryer, a typical value for the amount of air allowed per square foot, ,
was assumed to be 750
ℎ ∗
using equation (65) below.
. With this information, the required diameter can be calculated
=
The diameter was found to be .
∙
̇
, which was rounded to
typical diameters are between 4 and 10
(65)
. The heuristic states that
so this seems reasonable. The residence time, �, was
assumed to be 1 hour, but further pilot plant testing would be required to accurately determine
the drying time for the specific feed used in this process.
The volume, , is now calculated using equation (66). A typical value for the area occupied by
solids,
, is assumed to be 8%. The bulk density of corn stover, , is reported as .
(Sudhagar Mani, 2004).
101
The volume required was calculated to be .
=
τ
̇
(66)
. From the diameter and volume, the
length required is determined from equation (67).
=
The length was calculated to be
(67)
. The maximum length of a single unit is determined
from heuristics which states that the lengths are typically between 4 and 15 diameters. The
diameter is
, which corresponds to a maximum length of
. Because of this, 4 units of
will be used in series to complete the drying process. The number of flights is
determined via heuristics and is between 2 and 4 times the diameter. For this design, that is 27
flights. The last important consideration is the RPM of the dryer, which is also calculated via
heuristics. This heuristic states the product of the diameter and RPM will be between 25 and
35. Thus, a value of 3.5 RPM was used, which satisfies the heuristic.
The construction material of choice for these rotary dryers would be carbon steel as the
biomass and air are not corrosive or dangerous and carbon steel is the most economic choice. A
summary of all the design specifications is shown below in Table 55.
Table 55: Design Summary for P-2 / RDR-101 Dryer
Design Specification
Number of Units
Length per Unit
Number of Flights per Unit
Diameter
RPM
Mass Flowrate
Material
Residence time
Value
.
.
Units
�
�
^
�
�
ℎ
�
ℎ
102
P-3 / GR-101 Grinding
The third step in the biomass pre-processing section is the stage 2 size reduction. This is a
necessary step to achieve the particle size distribution necessary to enter the gasification
reactor and is specified in Table 56.
Table 56: Size Range for Particles Leaving Stage 2 Grinding
Size Interval Lower Limit (�
Size Interval Upper Limit (� )
Weight Fraction Biomass Particles
100
120
0
120
140
.1
140
160
.2
160
180
.3
180
200
.4
A schematic of a typical hammermill is shown below in Figure 13 (Feed Machinery, 2015).
Figure 22: Hammermill General Design
The stage 2 grinder is modeled as a hammermill and has a power requirement of
.
ℎ
. The feed is delivered to the hammermill via a rolling conveyer but the outlet of
the feed must be changed into a closed-vessel pneumatic conveying system that prevents
entrainment of the new micron-sized particles. The operating conditions are displayed below in
Table 57.
103
Table 57: Operating Conditions for P-3 / GR-101 Grinder
Design Specification
Power Requirement
Value
1.14×106 Btu/hr
Capacity
2.04×104 lb/hr
Pressure
1.013 bar
Temperature
90°C
Important considerations for designing a hammer mill include the power requirement required
to drive the machine, the capacity, materials of construction, size of the mesh screen,
dimensions, and RPM. The Bliss Hammermill EMF was decided upon due its previous use in
processing sorghum stover (Neal A. Yancey, 2011) as well as it meeting the design specifications
above.
The Bliss Ha
e
ill utilizes utte a d/o i pa t plates f o
i e o lo k to th ee o lo k a d
is capable of grinding to finer particle sizes than traditional hammermills. An extra fine grind
screen is an addable option which prevents leakage. This is recommended as the particles will
be between 100 and 200 microns so leakage is a concern. The Hammermill specifications are
shown below (Bliss Industries, Inc.) in Table 58.
Table 58: Design Specifications for Bliss Hammermill EMF 4848
Model #
4848
Power (Btu/hr)
1.02×106 to 1.27×106
Diameter (in)
48
Width (in)
48
RPM
1800
This model also includes a dual grind chamber and sliding doors as standard equipment. From
this information, the power requirement is met and is assumed that the capacity is met with
one unit as power is a more useful metric to judge the capability of the machine. The screen
size required to meet the particle size distribution is 100 MESH which has a max opening of 150
microns. The material of construction is carbon steel with replaceable wear parts. This
information is compiled below in Table 59.
104
Table 59: Summary of Design Considerations for P-3 / GR-101 Grinding
Design Specification
Model
Power
Diameter
Width
RPM
Capacity
Material of Construction
Screen Size
Value
4848
1.14×106
48
48
1800
2.04×104
Carbon Steel
MESH 100
Unit
Unitless
Btu/hr
in.
in.
revolutions/minute
lb/hr
Unitless
Unitless
P-4 / HP-101 Hopper
The last unit operation in the biomass pre-processing section is a lock-hopper. The role of the
lock-hopper is to pressurize the now micron-sized particles from .
to
.
necessary for the gasification reactor and is accomplished by feeding nitrogen gas at
. This is
.
.
The flowrate of nitrogen required to pressurize the feed stream is calculated in the simulation
to be .
ℎ
.
A lock-hopper is two hoppers connected via a sealable feeder system. The top hopper contains
a bin that is able to hold a buildup of biomass in case of any plant shutdowns or delays. This is
operated in a semi-continuous fashion and the feed is fed into the lower hopper, which seals
and pressurized gas is flowed in. Once pressurized, the bottom hopper is allowed to continue to
the gasification reactor and the process is repeated. Pneumatic conveying is used to both feed
the lock-hopper as well as to carry the now-pressurized biomass to the top of the heliostat
where the gasification reactor resides. An image of a general lock-hopper is shown below in
Figure 14 (Chemical Processing, 2006) and an image of what a typical exit utilizing pneumatic
conveying is shown in Figure 15 (Steam of Boiler, 2015).
105
Figure 23: General Lock-Hopper Design
Figure 24: General Outflow from Rotary Feed to Pneumatic Conveyance
The important design considerations for a lock hopper are the angle of the wall, bin volume,
mass flowrate, power consumption for agitation, and the size of the hopper opening. To design
the lock hopper, the method outlined in Introduction to Particle Technology is used (Rhodes,
2008). To determine the necessary design specifications, the angle of internal friction as well as
the effective angle of internal friction is required. These require physical lab-scale testing yet a
106
report on corn stover that is larger than our particles was found and assumed to apply to our
particles (Nehru Chevanan, 2009). Their research reported the following values, which can be
found in Table 60.
Table 60: Reported Results from Corn Stover Hopper Flowability Parameter Research
Parameter
Effective Angle of Internal Friction
Angle of Internal Friction
Value (Degrees)
Symbol
�
Using this information and Figure 16 which corresponds to a conical hopper with an effective
angle of internal friction, , of 60 degrees, the wall angle is able to be determined.
Figure 25: Figure corresponding to effective internal angle of friction of 60 degrees. Used to determine wall angle.
From this, the internal angle of wall friction, � , is not on the figure but is actually on the top
left of this diagram. This points to a hopper wall angle of 0 degrees. It is noted in the paper that
107
corn stover of their size range is probably not suited for hopper design and might not flow
under gravity alone. Their particle size range is significantly larger from the biomass in the
simulation and it is likely that due to the smaller nature of the particles that they will flow more
easily under gravity. Noting the discrepancy and without further knowledge of more specific lab
data, the assumption will be made that the system can be operated with a lock-hopper with an
angle of 0 degrees.
To determine the design specifications, Equations (68) through (72) along with the table of
nomenclature in Table 61 below are used.
Table 61: Table of Nomenclature for P-4 / HP-101 Hopper
Symbol
Description
Velocity
̇
Mass flowrate
̇
Volumetric flowrate
�
Hopper Cross-Sectional Area
Hopper Height
Biomass bulk density
�
Hopper Volume
Time for hopper storage
�
Using a length of
=
ℎ
(68)
from the dryer design process, the velocity was found to be .
Using a biomass mass flowrate of .
was calculated to be .
Unit
.
̇ =
̇
(69)
with a bulk density of .
�
=
.
̇
, the volumetric flowrate
(70)
108
The minimum area for the hopper opening, using the above volumetric flowrate and velocity,
was calculated to be .
. To construct a reasonable height for the silo, along with the
process operating semi-continuously with a controlled feeder, the area was enlarged by a factor
.
of 2. This new area for the hopper opening is
�
=
.
̇
∗
(71)
It was assumed that 20 minutes of storage capability was adequate in case of a plant shutdown
to allow for the feed to stop. Using the volumetric flowrate and that time, the volume of the
.
top hopper, which includes the storage silo, was found to be
. The bottom hopper
was assumed to be small so that loading times are short and the pressurization process has a
small residence time. It was decided to use 1 minute of storage for the bottom hopper. Using
these values, it was found that the volume required is
�
=
�
�
.
The height of the top hopper with silo was calculated to be
height was calculated to be .
.
.
(72)
.
. The otto
hoppe s
A rotary feeder system is used in the lock hopper valve to control the flowrate as well as seal
the pressurization hopper. Another rotary feeder hopper is located under the bottom hopper to
feed into the pneumatic conveying system as shown in Figure 15 above. It is assumed that 5%
of the nitrogen used to pressurize the biomass is leaked into the gasifier and this is modeled in
the simulation as the component splitter. It is assumed that the material of construction is
carbon steel as it economic and there are no special considerations needed for the solid
biomass feed. The power consumption used for agitation is found from the simulation to be
.
ℎ
. Table 62 below summarizes the design of the lock-hopper.
Design Specification
Hopper Wall Angle
Hopper Area of Opening
Top Hopper/Bin Volume
Bottom Hopper Volume
Table 62: Design Summary for P-4 / HP-101 Hopper
Value
Units
.
818.4
40.92
109
.
Power Consumption
Velocity
.
Biomass Mass Flowrate
.
.
Nitrogen Mass Flowrate
Solar Field and Tower
ℎ
ℎ
ℎ
In order to achieve the high temperature requirements of the gasification reactor, and to lower
GHG emissions from the plant operation, concentrated solar power is utilized to operate the
reactor for 8 hours each day. Many factors are considered when designing a solar heliostat
field. Initially, the solar radiative power is taken into consideration. The rate at which solar
energy in the form of electromagnetic radiation reaches the earth is solar irradiance and is
measured as power per unit area, W/m2 (Power from the Sun, n.d.). This irradiance varies with
time and location, but when designing a system the maximum solar irradiance value is used to
determine the peak rate of energy input into the system (Power from the Sun, n.d.). For the
design calculations, the maximum solar irradiance was assumed to be 1000 W/m 2. Realistically,
this value undergoes cosine losses, which is the loss of radiation when the solar energy hits the
earth at any angle outside of normal to the Sun.
The heat duty from the gasification reactors was used to determine the size of the tower and
heliostat field. For a heat duty of 6.5e4 kWh/hr, the net energy required to the process is 521.9
GWhr. The solar potential is 261 GWhr, though solar only delivers 205.6 GWhr to the process
while natural gas supplements the remaining 316.3 GWhr. The excess is due to variations
e pe ie ed o e ti e i sola e e g a d is t accounted for in this design project. A
recommendation to mitigate the
asted energy could be to stop tracking some of the
heliostats during times of excess solar.
Using the required reactor duty, a tower height and heliostat area was calculated at 4000x and
8000x concentrations. The goal is to minimize costs between the number of towers and
heliostat area, while still providing enough power to the process with acceptable efficiencies. A
calculation with 8000x concentration provided a tower height outside of the limit of 200 m.
When analyzing the 4000x concentration design, a single tower height of 182 m fit within heat
110
duty specifications. The material of choice would be a cement composite constructed to
withstand the weight of the interior reactor and heat requirements from the field.
Analyzing the 4000x concentration parameters to find the optimal design involves evaluating
tradeoffs between acceptance ( ) and lookout angle ( ), field size, CPC size, heat losses, and
efficiency. As the acceptance and lookout angles increase, the size of the field increases, which
allows for more power to the CPC. The trade-offs to this design are larger heat losses and larger
CPC area, which affects efficiency and cost. The key design point to consider is at day 82 hour
10, as this point mimics average conditions as closely as possible. For an acceptance and
lookout angle of 35, the design point parameters are listed below in Table 63.
Table 63: Key design parameters for solar field
ηtp
ηs
95%
Power,
Heliostat
area, m
2
Total field
area, m
2
Total CPC
area, m
2
# of
Heliostat
heliostats
dimensions, m2
2802
9.1x9.1
kW
0.751
0.649
6.64x104
2.320x105
2.326x105
302
The heliostats are single elements with square geometry and a spherical curvature. Their
dimensions are 9.1m x 9.1m. They are designed to be highly reflective mirrors with low optical
errors. In the field, they are closely packed and slightly offset to minimize shading and blocking
losses. This can be seen in the extremely small difference between field area and total heliostat
area. The heliostats will be oriented in three fields around the tower at 120 angles from each
other. Total land area was calculated by determining the land requirements for the north field
and assuming the other two fields have the same performance specifications. In reality, each
field would have slightly different size requirements and efficiency outputs.
Three compound parabolic concentrators (CPC) are oriented near the top of the tower with an
entrance aperture directly facing each field. As an estimation, it is assumed the CPC can
generate 95% of power, which allows for design optimization without using reactor
temperature as a factor (Lewandowski). The entrance radius is larger than the exit to provide
the required concentration of thermal energy to meet the power and exiting heat flux
requirements of the reactor. This optimization was calculated by the design spreadsheet so the
111
aperture area was minimized, which minimizes reradiation heat losses within the reactor tubes.
The overall results are the CPCs require 302 m2 of surface area with an entrance radius of 3.7m,
exit radius of 2.15 m, and height of 6.73 m. Figure 26 shows an example of the orientation of
the solar field and heliostats.
Figure 26: Solar field orientation showing heliostat and CPC layout
Based on the specified design point, the efficiency to the process,
, is 0.751. This value
reflects capturing all of the radiation reflected by the heliostat field, including solar efficiency
with reflective losses and shading/blocking losses from the heliostats. The overall solar
efficiency is the amount of instantaneous power into the reactor, including losses from the sun
to the reactor via the heliostats (Lewandowski). At the design point averaged for all three fields
the overall solar efficiency,
, was determined to be 0.60. The overall yearly solar efficiency
averaged for all three fields was determined to be 0.51.
Biomass Gasification
Solar Reactor
In this process, the solar reactor serves three critical functions: (1) to burn the biomass broken
down by the preprocessing section to char (carbon), (2) facilitate the char-gasification reactions
to create syngas composed of carbon monoxide (CO) and hydrogen (H2), and (3) optimize the
H2:CO ratio using steam-methane reformation. In sizing the reactor, the kinetics of the system
first had to be modeled. In the solar reactor, there are many reactions occurring
simultaneously, but only the char gasification reaction shown in Equation (73) was assumed to
112
be rate-limiting. To be rigorous, however, additional reactions were included in the model and
are presented in Equations (74) to (77), which ultimately led to a 0.4 m3 addition of volume to
the reactor compared to when only the char-gasification reaction was considered. The steam
methane reforming reactions (SRM1 and SRM2) are shown in Equations (74) and (75), the
water gas shift (WGS) reaction is shown in Equation (76), and the dry reforming of methane
(DRM) reaction is shown in Equation (77).
+
→
+
(73)
+
(75)
+
(77)
+
↔
+
(74)
+
↔
+
(76)
+
↔
+
↔
Once the important reactions were identified, the kinetic rate for each reaction was found in
literature. The kinetic expression suggested by R. Bryan Woodruff and Alan Weimer for
modeling char gasification is a random pore plugging model and is presented in Equation (78)
(Woodruff & Weimer, 2013). The kinetic constants are summarized in Equation (79) and Table
64.
ℎ
=
=
+
−
√ −�
+
−
(78)
(79)
Table 64: Summary of constants used in kinetic model for solar reactor, where
Equilibrium
constant, ��
�
=
�
Pre-exponential
factor, ��
2.51×103
Activation Energy,
� (kJ/mol)
112.6
Units for
�� , �
6.74×10-2
-37.3
bar-1
3.04×10-1
-36.6
bar-1
− �/
bar-1s-1
Here, conversion, , is defined as moles of product gas over moles of total gas, using Equation
(73) to define product gas. Since in an ideal system, moles are essentially directly proportional
to pressures, an approximation of the component mole fraction
partial pressures of gases in the mixture
�,
using Equation (80).
�
was used to determine
113
where
�
=
�
∙
(80)
is the total pressure of the system. The paper introduced a geometry factor, �, and
discussed its dependence on reaction conditions. Since no correlation for determining the value
of this parameter was presented based on temperature or pressure, the same value of 4.3
mentioned in the paper was used. Pilot plant testing would be required to determine � more
accurately. It is also clear that pilot plant testing will be required in order to determine the
accuracy of the utilized kinetic model at the conditions of plant operation.
In modeling the SRM1, SRM2, WGS, and DRM reactions, no models were found for the uncatalyzed reactions. To ignore these reactions, however, would lead to an underestimation of
the required reactor volume. To compensate, a catalyzed, competitive binding reaction model
proposed by Jun, et al. was used and is presented in Equations (81) to (84), even though the
system used in the provided paper did not exactly match that used in the proposed process
(Jun, et al., 2011). The kinetic parameters are summarized in Table 65. By using just 1 g of
catalyst, it was proposed that this would be an approximation of un-catalyzed conditions.
=
=
=
[ +
+
[ +
+
[ +
+
=
( +
+
+
(
+
+
−
+
−
+
−
/
/
)/
−
+
)( +
.
(81)
.
(
)]
(82)
(
)]
(83)
(
)
)]
(84)
114
Table 65: Parameters for SRM1, SRM2, WGS, and DRM reaction rates (Jun, et al., 2011)
Other undesired side reactions that were not specifically modeled include creation of nitrogen,
amine, hydrogen chloride, ethane, ethene, nitric oxide, sulfur, and combustion products of
methanol. A linear rate was used to account for their presence in the total moles over the
reaction progress, which approximated their effect on the other rate expressions used.
The kinetics were modeled using Excel, as presented in Appendix II-L and residence time, �, was
determined. The conversion of carbon was relative to a feed of 110% of the fed carbon in order
to account for the recycle stream entering the solar reactor. This residence time was then used
in conjunction with a geometric average between the entering streams and exiting stream
115
volumetric flow rates, , to determine the reactor volume, , that would be required to reach
the 98.18% conversion of char specified by the Aspen PLUS simulation, as shown in Equation
(85). The geometric average was selected as an appropriate approximation of the average
volumetric flow rate flowing through the reactor, since conversion happens more rapidly near
the beginning of the reactor than near the end. The results of conversion versus residence time
are summarized in Figure 27.
=
1.0
∙�
(85)
0.9
0.8
Conversion, X
0.7
0.6
0.5
0.4
0.3
0.2
0.1
0.0
0
1
2
3
4
Residence time, τ (s)
5
6
Figure 27. Conversion of carbon versus time spent in the solar reactor. Note that conversion at the inlet is already at about 50%
due to gas present from the preprocessing section and from the recycle stream.
The estimation of required reactor volume was 6.58 m3. In order to withstand the reaction
te pe atu e of
ha e a i
˚C, sili o
e dia ete of
a ide “iC tu es e e e ui ed. Since the available SiC tubes
a da e
lo g, it was determined that 18 tubes would be
required to achieve the desired conversion of carbon. These tubes would need a surrounding
p essu e essel o posed of stai less steel ith a
-thick alumina coating to protect the
vessel from the heat and to insulate the cavity. The required heating duty is supplied by natural
116
gas for an estimated average of sixteen hours per day, thus burners are inserted into the
bottom of the cavity. For the other eight hours of the average day, the heating duty is supplied
by the solar field, thus three apertures are included in the side of the pressure vessel to allow
concentrated sunlight from the solar field in. The tubes would be aligned in staggered
formations of three and separated around the three apertures, as shown in Figure 28, and the
apertures would be aligned with the CPCs described in the previous section.
Figure 28. Solar reactor drawing. The tubes create a fully defined diameter of 112" by separating the circle into 3 equal portions
with a lateral le gth e ual to twi e the width of the alu i a oati g plus the le gth of 6 adja e t 6 dia ete , / thi k “iC
tubes. Then, determining the diameter using Pythagorean Theorem yields D2=L2+(L/2)2, or D=L/2*(5)^0.5.
In designing the geometry of the stainless steel pressure vessel (metal casing), the tubes were
used as the li iti g fa to .
dia ete
ith ¾ thi k ess, a spa i g of ¼ ,
thi k oati g of
alumina, and three equivalent sections of 6 tubes each fully defined the diameter of the vessel
by Pythagorean Theorem, as described by Equations (86) and (87). L is defined as the length
117
o upied
the tu es plus the le gth o upied
the alu i a oati g
of spa i g. This esulted i a et dia ete of
plus a s all a ou t
. The height was taken to be a little higher
.
than the height of each tube, so about
=
+( )
= √
(86)
(87)
Finally, the biomass being fed to the system from the preprocessing section previously defined
contains a mixture of fluids and solids that must be properly fed to the reactor. To do this,
pneumatic conveying would be used, as previously described in the preprocessing section.
Essentially, a rotary screw would be used to control the flowrate of gas concurrently with
entrained solids.
Zinc-Oxide Reactor
The zinc-oxide (ZnO) reactor serves to remove most residual acids from the system before
moving on to the methanol reactor, where these acid gases could poison the catalyst. ZnO can
be used to remove both HCl and H2S from the syngas, but essentially no HCl is present in the
system after the amine scrubbing system, so the only kinetics requiring modeling are those of
Equation (88).
+
↔
+
(88)
Huiling, Yanxu, Chunhu, Hanxian, and Kechang studied the adsorption of sulfur by ZnO in the
presence of hydrogen and found that hydrogen had a significant effect on the rate of hydrogen
sulfide adsorption. The study suggested the use of an equivalent grain model presented in
Equation (89) (Huiling, Yanxu, Chunhu, Hanxian, & Kechang, 2002). The necessary parameters
are summarized in Table 66.
[ � ]=
+
(89)
118
Table 66: Summary of kinetic parameters used to determine residence time in the zinc-oxide reactor
Parameter
Equation/Value
Definition
Constant based on grain
Constant based on pellet
[mol/cm2 min]
[mol/cm2 min]
−
.
.
−
×
×
−
−
−
−
/
−
/
ex� −
Grain conversion function
−
ex� −
Pellet conversion function
Apparent chemical reaction constant
Diffusion coefficient
0.99
Conversion
[mol/L]
20.37
ZnO concentration in pellet
[nm]
26.1
Grain radium
[g/Nm3]
0.032587133
Pellet radium
[J/K mol]
8.314
Gas constant
[K]
483.15
Temperature
The time determined from Equation (89) was used as the residence time in Equation (85) to
determine the reactor volume necessary to achieve 99% conversion of H2S. Once again, the
reactor volume was sufficient to determine the length and diameter of the pressure vessel.
Monel-400 was chosen as the material of construction as there was sufficient acid present in
the stream to be corrosive. The required catalyst was then determined from stoichiometry and
the conversion percentage.
Methanol Reactor
The methanol reactor serves to produce the methanol product from syngas. The feed to the
reactor contains a H2:CO molar ratio of about 2 and a H2O:CO molar ratio of about 0.1, which
was specified as the optimal conditions for the reaction in the project description. The kinetic
model suggested by Bussche and Froment involved a two-step mechanism described by
Equation (90) (Bussche & Froment, 1996).
+
↔
+
+
↔
+
(90)
119
Bussche and Froment also suggested the use of a series of reversible binding reactions to the
catalyst as the kinetic model for the reverse water gas shift (RWGS) and methanol synthesis
reactions. However, Vijayaraghava and Lee suggested a similar model with a better fit for the
methanol synthesis step (Vijayaraghava & Lee, 1993). In order to get the best model of the
system kinetics, a hybrid of the RWGS model from Bussche and Froment and the methanol
synthesis model from Vijayaraghava and Lee was used. These models are presented in
Equations (91) and (92), respectively. The parameters of for these rates are summarized in
Table 67 and 68, respectively.
[
[
∙ℎ
∙
]=
]=
+
+(
∙ ex� −
′
∙
)(
+
∙
∙
∗
[ −
∙
−
+
(
(91)
∙
)]
)+√
(92)
+
Table 67: Summary of parameters for methanol formation reaction
Parameter
Equation or Value
[Cal/mol]
[MPa (kmol solute/kmol solvent)]
[MPa (kmol solute/kmol solvent)]
[MPa (kmol solute/kmol solvent)]
ex� (−
ex� (
.
ex�
− .
.
+
−
log
18360
.
( −
−
.
+ .
.
.
+ .
−
ln )
) / .
−
×
965.96
.
ln )
+
.
.
0.0150
1.49×10-3
3.96×10-3
3.68×10-5
0.818
0.823
2.090
2.160
120
Table 68: Summary of parameters for the WGS reaction
Parameter
√
′
Pre-exponential Factor
Activation Energy or (-��) [J/K-mol]
0.499
17197
6.62E-11
124119
3453.38
--
12200000000
-94765
This kinetic system was modeled using Polymath, the code for which can be found in Appendix
II-L. A graph depicting conversion versus residence time can be found in Figure 29. The
residence time required to reach 45% conversion was determined to be 0.02 hrs, and Equation
(85) was used again to determine the reactor volume. Since the reactor volume was 1094 m3
(much larger than the maximum purchasable reactor listed in the design textbook), it was
decided that the reactor would be split into three equivalent reactors. The exact design
specifications can be found in Appendix II-L. The chosen material of construction was 316
stainless steel since it would be capable of withstanding the large pressure requirements of the
Conversion
vessel.
0.5
0.45
0.4
0.35
0.3
0.25
0.2
0.15
0.1
0.05
0
0
0.005
0.01
0.015
Residence time (hr)
0.02
0.025
Figure 29. Residence versus residence time in methanol reactor
121
The heat of reaction for the formation of methanol is significantly negative, indicating that heat
exchangers would be required for each reactor in order to maintain the reaction condition of
210°C and avoid a runaway reaction.
One more important factor in designing the methanol reactor was the inclusion of the catalyst
slurry. Vijayaraghava and Lee described the reactor conditions as being composed of 25:1
volume fraction of slurry. The slurry was composed of 500 mL Witco-40 oil containing 25 g of
Cu/ZnO/Al2O3 catalyst. This, however, was not a sufficient catalyst content to keep the reactor
volume to a reasonable level, so the design was modified to include a much more concentrated
form of the slurry with 625 g of catalyst per 500 mL of Witco-40 oil. This ended up accounting
for most of the total reactor volume (25*1094 m3/26=1052) and was found to be a major
operating cost.
Putting all this information together led to the design of three vertical pressure vessels filled to
25/26ths of their volume with Witco-40 oil containing 625 g of Cu/ZnO/Al2O3 catalyst. The
syngas would then be bubbled through the large reactors, and the reactions would take place in
the slurry mode.
Spray Quench Tank
The spray quench tank used in the biomass gasification subsystem of the facility was modeled
as the reductant contact vessel present in a selective catalytic reduction (SCR) process.
In selective catalytic reduction, fly ash-containing flue gas from coal-fired power plants is
stripped of NOx gases by contacting the gas with a reductant such as ammonia or urea (Heck,
1999). A diagram of a simple SCR reductant contact vessel is provided in Figure 30.
122
Figure 30. Schematic of a simple selective catalytic reduction contact vessel.
The inspiration for this design choice was derived from the twin selective catalytic reduction
vessels used at Pawnee Station, a 500MW coal-fired power generation facility owned and
operated by Xcel Energy in Brush, CO. The vessel used at Pawnee Station is comprised of a
stainless steel shell approximately four inches thick, with a central spinner head made from
titanium. The reductant used at Pawnee Station is an amine-containing liquid, and is ejected
from the spinning head so that contact with the flue gas can be maximized. The selective
catalytic reduction system at Pawnee Station was custom-designed by Babcock and Wilcox
(B&W) for use as part of a greenhouse-gas emission reduction program.
In the variant of the SCR vessel modified to model the spray quench tank, the product stream
exiting the solar reactor enters the titanium rotating head, and contacts the cooling water
through a stainless steel shell. The original titanium material chosen for the rotating head by
Xcel Energy was selected in order to tolerate the corrosive environment created by the amine
fluid passing through it; for the purposes of the proposed process, the titanium serves as the
best material choice to withstand the high temperatures of the solar reactor product stream.
The stainless steel outer vessel casing is also an acceptable choice for the current application, as
123
the exiting cooling water stream leaves the process at high temperatures (586°C) and high
pressures (35 bar). Despite the hot solar reactor product stream contacting with the stainless
steel shell during heat transfer with the cooling water, substantial heat is lost when the stream
exits the rotating head nozzles. This heat loss is sufficient in order to prevent the stainless steel
shell from failing following high thermal contact (Alexander, 2015).
The SCR vessels in operation at Pawnee Station are approximately 20ft in diameter at the
widest part, and narrow to a 5ft diameter spent reductant chute. The vessels are approximately
80ft tall, measured from the base of the reductant chute to the top of the flue gas exit.
Heat Exchangers
Three heat exchangers were used in the Biomass Gasification subsection of the Solar-Thermal
Biomass Gasification facility designed here. The procedure and equations used to design each
heat exchanger has already been discussed in the General Equipment Design subsection,
however, metrics unique to each heat exchanger are presented below.
HEAT-1
The first heat exchanger present in the biomass gasification subsection, Heat-1, increases the
temperature of the clean gas (CLN-GAS) stream prior to entering the zinc reactor (ZN-REACT).
The heating fluid used for this heat exchanger is high pressure stream, which enters the process
at 450 psig and 231°C with a mass flowrate of 1.61x104 lb/hr. The utility fluid condenses as it
passes through the exchanger, supplying heat to the process stream. The process stream enters
the exchanger at 20.0°C and 35.0 bar, and exits the exchanger at 210°C. A summary of the heat
exchanger parameters are detailed in the following table.
Table 69. Heat-1 heat exchanger design parameters.
Heat duty
1.23×107 Btu/hr
Heat transfer area
1619 ft2
Tube O.D.
0.75 in.
Tube thickness
14 BWG
Tube length
16 ft.
Tube spacing
Triangular
Number of tubes
516
124
Number of shell and tube
One shell pass
passes
One tube pass
Pressure drop on tube side
3 psi
Pressure drop on shell side
1.5 psi
Materials of construction
Tubes: stainless steel
Shell: carbon steel
NH3-SEP HX
The second heat exchanger present in the biomass gasification subsystem is NH3-SEP HX, which
reduces the temperature of the vapor stream exiting the cyclone (VAPOR) prior to entering the
NH3-SEP unit. The cooling fluid used to achieve the desired temperature change is refrigerated
brine, which enters the separator at -17.8°C and exits the separator at 10.0°C. 2.99x105 lb/hr of
cooling fluid is required to achieve the desired heat transfer. The process stream enters the
heat exchanger at 122°C, and exits at 20.0°C. A summary of heat exchanger metrics are
provided in the following table.
Table 70. NH3-SEP HX heat exchanger design parameters.
Heat duty
9.16×106 Btu/hr
Heat transfer area
791 ft2
Tube O.D.
0.75 in.
Tube thickness
14 BWG
Tube length
16 ft.
Tube spacing
Triangular
Number of tubes
252
Number of shell and tube
One shell pass
passes
One tube pass
Pressure drop on tube side
3 psi
Pressure drop on shell side
5 psi
Materials of construction
Tubes: stainless steel
Shell: carbon steel
VL-SEP HX
The final heat exchanger present in the biomass gasification subsystem is VL-SEP HX, which
reduces the temperature of the product stream exiting the methanol reactor (RCT-PROD) prior
125
to entering the VL-SEP unit. This product stream enters the heat exchanger at 210°C, and exits
the heat exchanger at 50.0°C. In order to achieve this temperature difference, cooling water is
used as the utility stream. A flowrate of 1.53x106 lb/hr of cooling water at 32.2°C enters the
heat exchanger, and exits at 48.9°C following heat transfer from the process stream. In order to
accomplish the required heat transfer, two identical heat exchangers are run in parallel. A
summary of heat exchanger metrics for a single heat exchanger in the VL-SEP HX unit is
provided in the table below.
Table 71. VL-SEP HX heat exchanger design parameters.
Heat duty
4.59×107 Btu/hr
Heat transfer area
4041 ft2
Tube O.D.
0.75 in.
Tube thickness
14 BWG
Tube length
12 ft.
Tube spacing
Triangular
Number of tubes
858
Number of shell and tube
One shell pass
passes
One tube pass
Pressure drop on tube side
5 psi
Pressure drop on shell side
1.5 psi
Materials of construction
Tubes: stainless steel
Shell: carbon steel
Compressor
In the biomass gasification subsystem, a compressor is used to increase the pressure of the
inlet stream to the methanol reactor. This compressor takes an inlet stream with a mass
flowrate of 1.08x105 lb/hr at 133°C and 35.0 bar and increases the pressure of that stream to
80.0 bar. This action also increases the temperature of this stream to 219°C. In order to specify
the required design parameters for the compressor, the power consumed by the compressor,
Pactual, and efficiency must be calculated. From these parameters, the compressor type can then
be chosen.
126
The first step required for this calculation is the determination of adiabatic head, which is
accomplished by the following formula:
Where
= ′
�
(
�
[
)
−
− ]
is the ratio of the constant pressure heat capacity (
heat capacity (
(93)
) divided by the constant volume
). In Equation 93, ’ is the gas constant in units of J/(mol·K),
�
has units of K,
and pressures are in units of kPa. The units returned for adiabatic head are kJ/kg, but can be
readily converted to meters for the following calculation of adiabatic power, in which the
adiabatic head (m) is multiplied by the mass flowrate (kg/s):
=
∙ ̇
(94)
This equation returns the value for adiabatic power in kW.
The adiabatic power requirement calculated above is then divided by the actual power
consumed by the compressor, obtained from Aspen PLUS. The resulting value is the compressor
efficiency:
=
(95)
For proper compressor specification, actual power consumption is converted from kW to hp, so
that a suitable compressor type can be chosen. Based on the parameters discussed here,
compressor type was chosen using the following figure:
127
Figure 31. Compressor type selection based on adiabatic head and volumetric capacity; Peters et al.
A summary of the parameters relevant to the compressor design selected for this application
are summarized in the following table.
Table 72. Compressor design parameters.
Pactual
η
Brake hp
5519.66 kW
0.72
7401.97
Stages
Compressor Type
1
Centrifugal,
Single Stage
Pumps
Three pumps are present in the biomass gasification subsystem. A rigorous methodology
featuring appropriate equations has already been provided in the General Equipment Design
subsection. Metrics unique to each pump in the system are provided below.
128
PUMP-1
The first pump present in the biomass gasification subsystem elevates the pressure of the H 2O
inlet stream (H20-1) prior to entering the solar reactor. This stream has a flowrate of 1.92x104
lb/hr, and experiences a pressure increase from 1.01 bar to 35.0 bar. Pump design metrics are
provided in the table below.
Table 73. PUMP-1 design parameters.
RPM
1750
Casting
3x2x6
Efficiency
45%
Stages
6
Total Head (ft)
213
ft. head/stage
35.5
Total Power
6 hp
Power/stage
1 hp
Impeller Size
6 in
Duty
7.04x104 btu/hr
PUMP-2
The second pump in the biomass gasification subsystem elevates the pressure of an inlet
cooling water stream (H20-2) prior to entry into the spray quench tank. The inlet water stream
has a mass flowrate o 5.95x104, and must undergo a pressure change from 1.01 bar to 35.0 bar
prior to entry into the reactor. Pump design metrics are presented in the table below.
Table 74. PUMP-2 design parameters.
RPM
3500
Casting
3x2x6
Efficiency
55%
Stages
2
Total Head (ft)
213
ft. head/stage
107
Total Power
15 hp
Power/stage
7.5 hp
Impeller Size
6 in
Duty
1.57x105 btu/hr
129
PUMP-3
The final pump in the biomass gasification subsystem elevates the pressure of an inlet water
stream (H2O-3) from 1.01 bar to 35.0 bar prior to entry into the methanol reactor. This stream
has a mass flowrate of 5.60x103 lb/hr.
Table 75. PUMP-3 design parameters.
RPM
1750
Casting
3x2x6
Efficiency
20%
Stages
11
Total Head (ft)
495
ft. head/stage
45.0
Total Power
5.5 hp
Power/stage
0.5 hp
Impeller Size
6.5 in
Duty
6.47x105 btu/hr
Vapor-Liquid Separators
Two vapor-liquid separators were designed for the biomass gasification and methanol
production section of the proposed plant. The first of these separators, NH3-SEP, operates at
20°C and 35 bar and functions to separate ammonia and water from the syngas stream. The
second separator, VL-SEP, operates at 20°C and 35 bar and separates water and the newly
p odu ed
etha ol
a
etha ol f o
the p o ess gas. The desig
o side atio s fo ea h
of these separators are outlined below.
NH3-SEP
NH3-SEP was designed as a horizontal vapor-liquid separator using the procedure outlined in
the Horizontal Separators design section (see pg. 86). A horizontal orientation was used to
reduce the capital cost of the process and because the gas-liquid ratio in this unit was high. The
stream specifications taken from Aspen PLUS that were used in this design are presented in
Table 76 below. The liquid retention time was chosen to be
=
� because the system was
130
=
assumed to be non-foaming. A liquid droplet diameter of
�
was used to ensure
adequate separation of entrained liquid droplets from the gas phase.
Table 76: Aspen PLUS stream specifications used in the design of NH3-SEP
Parameter
Symbol
Value
Units
Gas compressibility
1.021
-
Evaluated at ACID-GAS
Gas flow rate
46.39
MMSCFD
Evaluated at ACID-GAS
Liquid flow rate
146.9
bbl/day
Evaluated at AQ-WASTE
Operating temperature
527.67
°R
Evaluated at outlet conditions
Operating pressure
507.6
psia
Evaluated at outlet conditions
150
µm
Minimum of 150-200 µm
Mass density of gas
0.951
lb/ft3
Evaluated at ACID-GAS
Mass density of liquid
62.33
lb/ft3
Evaluated at AQ-WASTE
0.015
cP
Evaluated at ACID-GAS
10.3
°API
Evaluated at AQ-WASTE
5
min
Table 47
Liquid droplet diameter
Viscosity of the gas
API gravity of liquid
Retention time of the liquid
°
�
From these specifications, a table of effective lengths
Comments
and standard internal diameters
was prepared according to Equations (46) and (47). The specifications that gave a slenderness
ratio
between 3 and 4 were chosen and used to calculate the seam-to-seam length
. The
results of these calculations are given in Table 77 below.
Table 77: Calculations of length and diameter for NH3-SEP
d (in.)
Leff (gas cap.)
Leff (liq. cap.)
Larger Leff (ft)
Sr
Lss (ft)
Lss < (4/3)Leff?
6
39.60
29.14
39.60
79.21
40.10
TRUE
12
19.80
7.29
19.80
19.80
20.80
TRUE
18
13.20
3.24
13.20
8.80
14.70
TRUE
24
9.90
1.82
9.90
4.95
11.90
TRUE
30
7.92
1.17
7.92
3.17
10.42
TRUE
36
6.60
0.81
6.60
2.20
9.60
FALSE
42
5.66
0.59
5.66
1.62
9.16
FALSE
48
4.95
0.46
4.95
1.24
8.95
FALSE
131
With these specifications and the properties of 316 stainless steel in Table 78 below, the shell
thickness was found to be
=
� ., which led to a vessel weight of
of the specifications of NH3-SEP are given in Table 79 below.
=
. A summary
Table 78: 316 stainless steel material properties (Nickel Development Institute), (Azo Materials, n.d.).
Metric
Symbol
Value
Units
Maximum allowable stress
19400
psi
Density
497.55
lb/ft3
Table 79: Specifications of NH3-SEP
Internal diameter
2.5 ft
Seam-to-seam length
10.42 ft
Material of construction
316 stainless steel
Shell thickness
3/4 in.
Vessel weight
3129 lb.
VL-SEP
VL-SEP was also designed as a horizontal vapor-liquid separator using the procedure outlined in
the Horizontal Separators design section (see pg. 86). A horizontal orientation was used to
reduce the capital cost of the process and because the gas-liquid ratio in this unit was high. The
stream specifications taken from Aspen PLUS that were used in this design are presented in
Table 80 below. The liquid retention time was chosen to be
of the liquid was greater than 35. A liquid droplet diameter of
=
=
� because the API gravity
�
was used to
ensure adequate separation of entrained liquid droplets from the gas phase.
Table 80: Aspen PLUS stream specifications used in the design of VL-SEP
Parameter
Symbol
Value
Units
Comments
Gas compressibility
1.019
-
Evaluated at SEP-VAP
Gas flow rate
50.41
MMSCFD
Evaluated at SEP-VAP
Liquid flow rate
4482.2
bbl/day
Operating temperature
581.67
°R
Evaluated at outlet conditions
Operating pressure
507.6
psia
Evaluated at outlet conditions
Evaluated at RAW-MEOH
132
Liquid droplet diameter
150
µm
Mass density of gas
0.906
lb/ft3
Evaluated at SEP-VAP
Mass density of liquid
48.52
lb/ft3
Evaluated at RAW-MEOH
0.016
cP
42.4
°API
Evaluated at RAW-MEOH
1
min
Table 47
Viscosity of the gas
°
API gravity of liquid
Retention time of the liquid
�
Minimum of 150-200 µm
Evaluated at SEP-VAP
From these specifications, a table of effective lengths
and standard internal diameters
was prepared according to Equations (46) and (47). The specifications that gave a slenderness
ratio
between 3 and 4 were chosen and used to calculate the seam-to-seam length
. The
results of these calculations are given in Table 81 below.
Table 81: Calculations of length and diameter for VL-SEP
d (in.)
Leff (gas cap.)
Leff (liq. cap.)
Larger Leff (ft)
Sr
Lss (ft)
Lss < (4/3)Leff?
6
57.52
177.87
177.87
355.73
178.37
TRUE
12
28.76
44.47
44.47
44.47
45.47
TRUE
18
19.17
19.76
19.76
13.18
21.26
TRUE
24
14.38
11.12
14.38
7.19
16.38
TRUE
30
11.50
7.11
11.50
4.60
14.00
TRUE
36
9.59
4.94
9.59
3.20
12.59
TRUE
42
8.22
3.63
8.22
2.35
11.72
FALSE
48
7.19
2.78
7.19
1.80
11.19
FALSE
With these specifications and the properties of 316 stainless steel in Table 78, the shell
thickness was found to be
=
� ., which led to a vessel weight of
of the specifications of VL-SEP are given in Table 82 below.
=
. A summary
Table 82: Specifications for VL-SEP
Internal diameter
3 ft
Seam-to-seam length
12.49 ft
Material of construction
316 stainless steel
Shell thickness
7/8 in.
Vessel weight
5394 lb.
133
ZN-SPLIT Cyclone
The ZN-SPLIT cyclone was used to extract the remaining zinc solid utilized in the zinc oxide
reactor. This cyclone design was similar to the cyclone design detailed in the General
Equipment Design of Cyclones. The inlet volumetric flowrate was higher than for the first
Cyclone unit operation, thus a larger diameter was required. Four units in parallel were
designed because the initial diameter calculation was outside of the range for a standard
diameter. The scaling factor was determined to be 2.18, which produced a final cyclone
diameter of 2.29 ft and height of 10.33 ft. ZN-SPLIT would not be in contact with any corrosive
material, therefore carbon steel can be used.
Amine Scrubbing
The design procedures used for the relevant equipment in the amine scrubbing section of the
plant are outlined in the following section. A summary of the designed equipment is shown in
Table 83 below. Procedures for some of the common equipment designed in this section are
given in Section VIII: Generalized Equipment Design.
Table 83: Summary of designed equipment for the amine scrubbing section
Piece of Equipment
Design details
Name in HYSYS Simulation
Absorption column
Absorber
Absorber
Vapor-liquid separator
Vertical separator
Separator
Lean/rich heat exchanger
Shell-and-tube heat exchanger
Lean/Rich Exchanger
Regeneration column
Distillation column
Regenerator
Regenerator condenser
Air-cooled heat exchanger
Regenerator
Regenerator reflux drum
Vertical pressure vessel
Regenerator
Regenerator reboiler
Kettle reboiler
Reboiler Steam
Booster pump
Centrifugal pump
Booster Pump
Amine cooler
Double-pipe heat exchanger
Cooler
Pressure Vessels, Separators, and Towers
The absorption column is the piece of equipment responsible for removing the acid gas
components from the process syngas, thereby sweetening it. To do so, intimate mixing
between the process gas and the amine solution is required. An absorption column with a set
134
number of sieve trays was designed to ensure this mixing. The liquid flows over a downcomer
on the trays, while the gas bubbles up through the gas and creates a froth. The height of liquid
volume must be large enough to ensure that there exists enough clearance for the liquid to
occupy. The absorption column itself was designed as a vertical vapor-liquid separator, as
explained in the Vertical Separators section (pg. 87); the height of liquid volume obtained from
this design procedure was used to inform the spacing of the trays in the column. The diameter
of the column was also obtained in this way. The shell thickness was obtained through
calculations demonstrated in the Pressure Vessels and Towers design section (pg. 90). The
following data in Table 84 from HYSYS and Arnold and Stewart in Surface Production Operations
were used in design calculations (Arnold & Stewart, 1999).
Table 84: Values for the design of the absorption column
Parameter
Symbol
Value
Units
Gas compressibility
1.014
-
Gas flow rate
46.265
MMSCFD
Liquid flow rate
759.7
bbl/day
Operating temperature
535.32
°R
Operating pressure
507.6
psia
Evaluated at outlet conditions
175
µm
Average of 150-200 µm
Mass density of gas
0.9574
lb/ft3
Evaluated at outlet conditions
Mass density of liquid
72.22
lb/ft3
Evaluated at outlet conditions
0.015
cP
10.1
°API
Average of rich and lean amine.
8
min
Table 47
Liquid droplet diameter
Viscosity of the gas
API gravity of liquid
Retention time of the liquid
°
�
Comments
Evaluated at outlet conditions
Average of feed gas and sweet gas
Evaluated at outlet conditions
Average of feed gas and sweet gas
Average of feed gas and sweet gas
Implementing the design procedure yields the specifications outlined in Table 85. The internal
diameter of the column was found to be
= .
; this value was rounded up to 4. The
value for height of liquid volume ℎ was found to be just under 24 inches. 24 inches is a standard
separation distance for trays according to Arnold and Stewart, so ℎ =
tray spacing.
=
� . was used as the
sieve trays were implemented, as recommended by Arnold and Stewart.
The seam-to-seam length
of the column was then computed by Equation (96):
135
.
=
∙ℎ
(96)
where the 1.15 figure accounts for a 15% disengagement length and the 12 figure accounts for
the conversion to feet. It was found that
=
. 316 stainless steel was used as the
material of construction for the trays and shell of the column to resist corrosive hydrochloric
acid. Using the properties of 316 stainless steel in Table 78, the shell thickness was found to be
=
� ., which led to a vessel weight of
=
.
Table 85: Specifications for the absorber column
Internal diameter
4 ft
Tray spacing
24 in.
Number of trays
20
Type of trays
Sieve trays
Seam-to-seam length
46 ft
Material of construction
316 stainless steel
Shell thickness
1 1/8 in.
Vessel weight
29515 lb.
The valve following the absorber column was not designed because valves are commonly
available from industrial vendor. In addition, additional information about the piping it would
be installed in would be needed.
The separator was designed as a vertical vapor-liquid separator. A vertical orientation was
hose
e ause of the u it s lo gas-liquid ratio. The same procedure was used to design this
separator as with the absorber column, with a few extra considerations. The vessel internal
diameter was calculated to be 0.538 ft; this value was rounded up to
achieve a slenderness ratio
= .
in order to
between 3 and 4. The shell thickness was required to be 1/4 in.
thick for structural stability. The material of construction was chosen to be 316 stainless steel to
resist corrosion. Specifications for the separator are shown in Table 86 below.
136
Table 86: Specifications for the separator
Internal diameter
2 ft
Seam-to-seam length
6.56 ft
Material of construction
316 stainless steel
Shell thickness
1/4 in.
Vessel weight
537 lb.
The regenerator column was designed as a distillation column. A common method used to
perform a preliminary design is the Fenske-Underwood-Gilliand (FUG) method, as was used to
design the methanol purification distillation column (pg. 145). However, this method breaks
down in the amine scrubbing system because of the electrolyte chemistry and non-idealities
present. This occurs despite the large boiling point difference (at 1 atm) between MDEA (
.
° ) and water (
ratio was calculated to be
=
�
=
° ), the heavy and light keys, respectively. The minimum reflux
= − . . So, suggestions from external sources were used to
guide the column design in HYSYS. The column was operated at total reflux since the product
gas must be treated and flared. 10 trays were chosen rather than 20 trays in this MDEA system
because the percent recovery of CO2 in the overall system is not increased enough to justify the
extra cost of nearly doubling the height of the column (Weiland & Sivasubramanian, 2003). The
reboiler duty was selected to be
= . ×
/ℎ to ensure that no minimum
temperature approach violations would occur in the Lean/Rich Exchanger and to reduce the
number of trays needed (Arnold & Stewart, 1999). An example calculation provided by
AspenTech was used to estimate the pressure drop of the column to be 4 psi (0.276 bar) and of
the condenser to be 2 psi (0.138 bar) (Aspen Technology, Inc., 2013). The condenser duty ( =
. ×
/ℎ ) and reflux ratio ( = .
were determined by the HYSYS simulation. The
distillate and bottoms flow rates, temperatures, and pressures can be found on Table 35. A tray
spacing of 24 in. was assumed.
For the calculation of the column diameter, however, the method outlined in the design of the
methanol purification column (see pg. 145) and Equation (104) was used. So, constant molar
overflow of the column was assumed. The shell thickness was found in the same way as a
137
pressure vessel. The material of construction was chosen to be 316 stainless steel. A summary
of important regenerator column specifications can be found in Table 87.
Table 87: Specifications of the regenerator column
Internal diameter
2.1 ft
Tray spacing
24 in.
Number of trays
10
Type of trays
Sieve trays
Seam-to-seam length
23 ft
Material of construction
316 stainless steel
Shell thickness
1/4 in.
Vessel weight
1834 lb.
Condenser duty
6.0×106 Btu/hr
Reboiler duty
4.7×106 Btu/hr
The regenerator was also equipped with a reflux drum to provide an avenue for liquid to flow
back into the column. Using the column environment from HYSYS for stream specifications, the
reflux drum was designed as a vertical separator. This decision was made because the gas-liquid
ratio was low. Once again, the internal diameter was rounded up to 2 ft in order to satisfy the
slenderness ratio requirement. The seam-to-seam length was also rounded up to 7 ft in order
to guarantee good liquid-vapor separation and reduce the loss of water droplets entrained in
the vapor. The results from these calculations are shown in Table 88.
Table 88: Specifications for the reflux drum
Internal diameter
2 ft
Seam-to-seam length
6.51 ft
Material of construction
316 stainless steel
Shell thickness
1/4 in.
Vessel weight
566 lb.
Heat Exchangers
The lean/rich heat exchanger was designed as a shell-and-tube heat exchanger using the
procedure explained in the Shell-and-Tube Heat Exchangers design section (see pg. 78) and the
information from Table 34. The two fluids exchanging heat are amine solutions, so an overall
138
heat transfer coefficient of
=
/ℎ ∙
∙ ° was chosen (Seider, Seader, Lewin, &
Widagdo, 2009). The choice of tube specifications was done to ensure the tube-side velocity in
the heat exchanger was between 1 and 10 ft/s. Stainless steel was used as the material of
construction to avoid the corrosive effects of hydrochloric acid. The final specifications are
summarized in Table 89 below.
Table 89: Specifications for Lean/Rich heat exchanger
Heat duty
1.45×106 Btu/hr
Heat transfer area
311 ft2
Shell O.D.
21.25 in.
Tube O.D.
1.25 in.
Tube thickness
14 BWG
Tube length
12 ft.
Tube pitch
Triangular
Number of tubes
80
Number of shell and tube
One shell pass
passes
One tube pass
Pressure drop on tube side
5 psi
(To Exchanger/Regen Feed)
Pressure drop on shell side
5 psi
(Regen Bottoms/To Tank)
Materials of construction
Tubes: stainless steel
Shell: stainless steel
The regenerator reboiler was designed as a shell-and-tube heat exchanger using the procedure
explained in the Shell-and-Tube Heat Exchangers design section (see pg. 78) and the
information from Table 36. One of the fluids in the heat exchanger was condensing steam, so
an overall heat transfer coefficient of
=
/ℎ ∙
∙ ° was chosen (Seider, Seader,
Lewin, & Widagdo, 2009). The log-mean temperature difference was assumed to be the
difference between the saturation temperatures of the tube and shell side fluids; this
assumption was made because one of the heat transfer media is changing phase. The choice of
tube specifications was made to ensure the tube-side velocity in the heat exchanger was
between 1 and 10 ft/s. Although not considered in the PFD description, estimates for the
139
pressure drop on the shell side and tube side were found using Heuristic 31. Stainless steel was
used as the material of construction to avoid the corrosive effects of hydrochloric acid. The final
specifications are summarized in Table 90 below.
Table 90: Specifications for the regenerator reboiler
Heat duty
��
6.00×106 Btu/hr
48.1°F (26.7°C)
Heat transfer area
384 ft2
Shell I.D.
21.25 in.
Tube O.D.
0.75 in.
Tube thickness
14 BWG
Tube length
12 ft.
Tube pitch
Triangular
Number of tubes
82
Number of shell and tube passes
One shell pass
One tube pass
Pressure drop on tube side
1.5 psi
(Saturated Steam/Steam Condensate)
Pressure drop on shell side
5 psi
(reboiler bottoms)
Materials of construction
Tubes: stainless steel
Shell: stainless steel
The other two heat exchangers used in the amine scrubbing system are not conventional shelland-tube heat exchangers. The first of these exchangers is an air-cooled heat exchanger for use
i the ege e ato
olu
s o de se . Design considerations were taken from Seide s
Heuristic 56 (Seider, Seader, Lewin, & Widagdo, 2009). The decision was made to use an aircooled heat exchanger to save on utility costs and to follow recommendations from Arnold and
Stewart (Arnold & Stewart, 1999). A schematic of one of these units is given in Figure 32 below.
An air-cooled heat exchanger utilizes ambient air to cool the process fluid in the tubes. The air
(cold stream) was assumed to enter at 90°F (32.2°C) and exit at 140°F (60°C). The inlet and
outlet conditions for the process fluid flowing through the tubes and the heat duty were
obtained from HYSYS. The process fluid (hot stream) enters at 235.94°F (113.3°C) and leaves at
140
221°F (105°C). Therefore, using Equation (37), the log-mean temperature difference is found to
be ∆
=
. °
. °
. An overall heat transfer coefficient of
for using unfinned tubes was assumed. From HYSYS, the duty was
amount of power needed for the fans was then found to be
= .
=
= . ×
/ℎ ∙
∙°
/ℎ . The
ℎ according to
Heuristic 56. From here, the same procedure to find the number of tubes in a shell-and-tube
heat exchanger was used. The final specifications are given in Table 91 below.
Figure 32: A general schematic of an air-cooled heat exchanger
Table 91: Final specifications for the air-cooled heat exchanger
Heat duty
-4.70×106 Btu/hr
Heat transfer area
3.30 ft2
Tube O.D.
3/4 in.
Tube thickness
14 BWG
Tube length
12 ft.
Tube spacing
Triangular
Number of tubes
1773 (unfinned)
Fan power required
1.34 hp
Materials of construction
Stainless steel tubes
A double-pipe heat exchanger was designed for use in the amine cooler. A double-pipe heat
exchanger was used here because the heat transfer area required in this process unit is less
141
than 100 ft2, making the use of a traditional shell-and-tube heat exchanger impractical. A
schematic of an example unit is given in Figure 33 below. The procedure used for calculating
the total heat transfer surface area was the same at that used for a shell-and-tube heat
exchanger. The utility used was chilled water; this fluid (cold stream) enters the shell at 45°F
(7.22°C) and leaves at 90°F (32.2°C). The lean amine solution enters the tube at 131.8°F (55.4°C)
and leaves at 110.0°F (43.3°C). The cooler duty is
transfer coefficient of
=
/ℎ ∙
= . ×
/ℎ . An overall heat
∙ ° for amine solutions and water was assumed
(Seider, Seader, Lewin, & Widagdo, 2009). The heat transfer area required was then calculated
to be
=
.
.
Figure 33: Schematic of a double-pipe heat exchanger
From here, two types of tubes were chosen – one was used as the outer tube, and the other as
the inner tube. The annular cross-sectional area
was then calculated according to Equation
(97):
where . .
=
. .
− . .�
is the inner diameter of the outer pipe and . .�
(97)
is the outer diameter of
the inner pipe. Using the heat capacity of water, the duty, and the surface area of the heat
exchanger, the mass flow rate of chilled water was then calculated to be
=
/ℎ . The
mass flow rate of chilled water was then divided by the density of water and the annular crosssectional area to find the superficial velocity of the chilled water to ensure that it was near the
142
range of 1-10 ft2. The inner tube material of construction was chosen to be stainless steel to
resist corrosion. The final specifications are presented in Table 92 below.
Table 92: Specifications of the amine cooler double-pipe heat exchanger
Inner tube
I.D.
3.068
in.
O.D.
3.5
in.
Schedule
40
Material of construction
Outer tube
Stainless steel
I.D.
4.026
in.
O.D.
4.5
in.
Schedule
40
Material of construction
Carbon steel
Heat duty
Q
1.9×105 Btu/hr
ft2
Heat transfer surface area
A
22.008
ft2
Mass flow rate of chilled water
m
4316.8
lb/hr
Tube side fluid:
Lean amine solution
Shell side fluid:
Chilled water
Pumps
The reflux pump, while shown in the PFD for the amine scrubbing system, was not given design
considerations and neglected. The design of the pump was neglected because the pressure
drop through the condenser and reflux drum was assumed to be negligible.
The booster pump was designed according to the pump design section (see pg. 75). The pump
curve used for the design is provided in Figure 34 below. From HYSYS, it was determined that
the pump needed to provide ∆ =
.
of head to
=
.
of lean amine flow.
HYSYS also provided an NPSHA of 42 ft. Final specifications for the booster pump are given in
Table 93 below.
143
Figure 34: Pump curves for the booster pump (Carver, 2006)
Table 93: Specifications for the booster pump
RPM:
Casing:
Efficiency:
Stages:
ft head/stage:
1750
3x2x6
37
rpm
(low flow)
%
6
38.93
ft
NPSHR:
5
ft
Power:
0.75
hp/stage
4.5
hp total
Impeller Size:
6.25
in
Product Separation & Post-Processing
Separator V-100
The separator in the methanol purification system served to create a 100% liquid inlet stream
to the distillation column and remove many of the gas impurities. A vertical separator was
designed for this process due to the low gas-to-liquid ratio of the inlet stream. This separator
was similar to the design calculations detailed in the Generalized Equipment Design: Vertical
Separators. The design diameter was calculated to be 30 inches and the seam-to-seam length
144
was 8.05 ft. Carbon steel is acceptable as the material for the separator because corrosive
compounds were not used in this unit.
Distillation Column T-100
A full distillation column was simulated in Aspen HYSYS with a total condenser and partial
reboiler. The choice for type and pressure of the condenser and reboiler was based on Figure
8.9 Algorithm for establishing distillation column pressure and condenser type in the class
Process Design textbook (Seider, Seader, Lewin, & Widagdo, 2009). A calculation was made for
the pressure of the distillate at 49°C; from this it was determined to use a total condenser at
2bar pressure and a reboiler at 2.758bar. The cooling fluid used for the condenser is chilled
water, which enters the process at 7.2°C with a mass flowrate of 7.6x104 lb/hr. The distillate
stream enters the exchanger at 85°C and 2.0 bar, and exits the exchanger at 80°C. A summary
of the heat exchanger parameters are detailed in the following table.
Table 94. Condenser heat exchanger design parameters.
Heat duty
3.74×107 Btu/hr
Heat transfer area
2225.4 ft2
Tube O.D.
1.5 in.
Tube thickness
14 BWG
Tube length
16 ft.
Tube spacing
Triangular
Number of tubes
143
Number of shell and tube
One shell pass
passes
One tube pass
Pressure drop on tube side
1.5 psi
Pressure drop on shell side
1.5 psi
Materials of construction
Tubes: stainless steel
Shell: stainless steel
The heating fluid used for the reboiler is low-pressure steam, which enters the process at
130.5°C with a mass flowrate of 4.21x104 lb/hr. The bottoms stream enters the reboiler at
129.2°C and 2.758 bar, and exits the exchanger at 129.8°C. A summary of the heat exchanger
parameters are detailed in the following table.
145
Table 95. Reboiler heat exchanger design parameters.
Heat duty
4.09×107 Btu/hr
Heat transfer area
915 ft2
Tube O.D.
1.25 in.
Tube thickness
14 BWG
Tube length
12 ft.
Tube spacing
Triangular
Number of tubes
234
Number of shell and tube
One shell pass
passes
One tube pass
Pressure drop on tube side
1.5 psi
Pressure drop on shell side
1.5 psi
Materials of construction
Tubes: stainless steel
Shell: stainless steel
An initial guess was made for the number of trays and entering feed tray, both to be modified
after hand calculations were performed. The Fenske-Underwood-Gilliland (FUG) shortcut
method was used to determine a variety of design factors for the distillation column. Equation
variables and nomenclature is given in Table 96.
Table 96: Table of Nomenclature for variables in distillation column design equations
Parameter
Symbol
Value
Units
variable
N/A
variable
N/A
N/A
N/A
Bottoms
N/A
N/A
Feed
N/A
N/A
calculated
N/A
calculated
N/A
N/A
N/A
N/A
N/A
0.466
N/A
0.7
N/A
Liquid mole fraction (x) for light key component (lk)
Liquid mole fraction (x) for heavy key component (lk)
Distillate
Relative volatility for lk and hk components
Initial mole fraction of component i in feed stream
ℎ
�
�
Value lies between the relative volatilities of the two key components
Moles of saturated liquid on the feed tray per mole of feed
Minimum reflux ratio at infinite L/D
Actual reflux ratio
∞
�
146
Actual number of equilibrium trays
23
N/A
11
N/A
34
N/A
1723.1
lbmol/hr
1395
lbmol/hr
Efficiency factor
0.7
N/A
Number of trays in the bottoms
28
N/A
Number of trays in the distillate
6
N/A
Fraction of vapor flooding velocity
0.8
N/A
Vapor flooding velocity
7.9
ft/s
0.2
N/A
0.001
lbmol/ft3
Liquid density
1.55
lbmol/ft3
Empirical capacity parameter
0.2
N/A
Parameter for sieve trays
0.2
ft/s
Surface tension factor
1.01
dyne/cm
Foam factor
1
N/A
Hole factor
1
N/A
Tower diameter
5.27
ft
Tower height
78.2
ft
Theoretical number of equilibrium trays
�
Actual number of trays
Total molar liquid flow rate
Distillate molar flow rate
Downcomer area/tower inside cross sectional area
⁄
Vapor density
First, liquid and gas mole fractions of each component in the condenser and reboiler were
determined, followed by a calculation of the k-values and relative volatilities of each
component. The Fenske equation uses this information to determine the minimum number of
equilibrium stages needed, as shown in Equation (98). This calculation determined the
minimum number of equilibrium trays to be 11.
�
=
ln[(
⁄
ℎ
) (
ln �
ℎ
ℎ
⁄
) ]
(98)
147
Next, an estimation was performed using the Underwood equations to obtain the minimum
reflux ratio,
∞
�
, to achieve ideal separation of the heavy key and light key components for
the process. Use of the Underwood equations involves two assumptions: first, that the liquid
molar flow in the rectifying section is constant, and second, that the relative volatility is
constant at the pinch point (Seider, Seader, Lewin, & Widagdo, 2009). The Underwood
equations are displayed below:
In order to solve for
∞
�
Σ
Σ
��,
��,
∞ �,
∞
��, ∞ �,
��, ∞ −�
=
−
=
, a guess is made for
until they come to unity. From this calculation,
+
∞
−
∞
(99)
(100)
�
and Equations (99) and (100) are iterated
�
was determined to be 0.4667 and the
actual reflux ratio to be 0.7. The Gilliland correlation estimates the actual number of
equilibrium stages,
, at a finite / for a specific ratio of ⁄
�
. This correlation has
limited accuracy in cases when the stripping section is more important than the rectifying
section in the overall separation. The correlation is most accurate for ⁄
and 1.5. Solving for
�
ratios between 1.0
in Equation (101) yields the actual number of equilibrium stages
necessary. Design calculations determined ⁄
−
+
�
= .
�
was 1.5, and
ln { − (
−
+
was 23.
�
)
.
}
(101)
An efficiency factor of 0.7 is used to calculate the actual number of design trays in the column,
which yielded 34 trays, as shown in Equation (102):
=
⁄
(102)
148
The Kirkbride equation was used for estimation of the feed stage in the distillation column.
allows back calculation to determine
Solving for
the value of
and
. The feed stage location would be
, rounded up to the nearest whole number.
ln (
)= .
ln {
ℎ
⁄
⁄
[
⁄
ℎ
] }
(103)
The tower diameter is dependent on the vapor and liquid flowrates, and is computed to avoid
entrainment flooding, which occurs when the liquid is entrained with the vapor due to too high
of vapor flow rates. The overall equation for a tower diameter calculation is shown in Equation
(104):
=[
The flooding velocity (
⁄
−
⁄
]
/
(104)
) is computed by Equation (105) shown below, where C is given in
Equation (106):
=
=
[
−
]
/
(105)
(106)
The diameter calculation is performed for both the bottoms and distillation section. The larger
diameter was chosen as the design diameter, which was 5.27ft from the rectifying section.
Tower height is calculated by assuming each sieve tray is 2 feet in height, and multiplying this
by
given in Equation (102). A 15% disengagement is added to this height to allow for
spacing at the top and bottom of the column. Tower height was calculated to be 78.2ft. As a
final step, the ratio of ⁄ should be calculated and was 14.82, which is reasonable as this
value should be kept less than 25. The
ate ial of hoi e is a o steel as the e a e t a
corrosive compounds flowing through the column.
149
IX. Utility Summary and Heat Integration
The entire process was considered for both heat integration and utility usage. Utility usage is a
very important factor when calculating operating expenses and heat integration is a very useful
tool that can be used to save considerable utilities by integrating heat already existing in the
facility to be used elsewhere.
Utility Summary
To obtain a closer look at how the entire process utilizes its utilities, an in-depth analysis was
completed broken down by individual unit operations that had a duty. A positive heat duty
implies energy is being added into the unit operation. Below in Table 97 is a list of the findings.
In-depth calculations can be found in Appendix IV-C: Summary of Utilities.
Table 97: Utility Summary Broken Down by Unit Operation
Equipment Identifier
P-1 / SR-101
Shredding
P-2 / RDR-101 Rotary
Drying
P-3 / GR-101 Grinder
Duty
(Btu/hr)
+ .
Type of Utility
Electricity
+ .
Electricity
Amine Reboiler
+ .
Amine Condenser
Blower
Amine Booster Pump
+ .
P-4 / HP-101 Hopper
Amine Cooler
MeOH Reboiler
MeOH Condenser
Pump-1
Solar Reactor
Solar Field
Methane Stream to
Solar Reactor
Consumption per Year
.
.
.
+ .
Electricity
Electricity
.
+ .
Medium Pressure
Steam
Electricity
.
+ .
Electricity
Chilled Water
.
+ .
Low Pressure
Steam
Chilled Water
.
+ .
Electricity
.
− .
− .
+ .
+ .
Solar
Natural Gas
.
.
.
Natural Gas
.
.
.
Cost per Year
($/yr)
.
.
.
.
.
.
.
.
.
.
.
.
.
.
150
Water Stream to
Solar Reactor
Quench Tank
Pump-2
NH3-SEP HX
HEAT1
COMPRESS
REACT
React Water
Pump-3
VL-SEP HX
XTRA-H2O
Cooling Water
− .
Cooling Water
+ .
− .
Refrigerant
.
+ .
High Pressure
Steam
Electricity
.
− .
Cooling Water
+ .
+ .
− .
Electricity
Cooling Water
Electricity
Cooling Water
Cooling Water
.
.
.
.
.
.
.
.
.
.
.
.
.
.
.
.
.
.
.
.
Shown below in Table 98 is the total usage and cost of each utility per year.
Table 98: Total Utility Usage and Cost per Year
Utility Type
Usage
Electricity
.
Chilled Water
Cooling Water
.
High Pressure Steam
.
Medium Pressure Steam
.
Low Pressure Steam
.
Natural Gas
Solar
Refrigerant
.
.
.
.
$
Cost ( )
��
.
.
.
.
.
.
.
.
.
151
From this information, the largest cost is the cooling water, which is the cheapest utility. The
next highest cost utilities are the natural gas, electricity, chilled water, and refrigerant. Using
this information, these would be the systems to analyze more clearly for gains in efficiency in
order to be more economically or environmentally efficient. The prices for each utility used to
calculate the above values is shown below in Table 99.
Table 99: Prices of Utilities
Utility Type
Cost
Electricity
Chilled Water
Cooling Water
High Pressure Steam
Medium Pressure Steam
Low Pressure Steam
Natural Gas
Solar
Refrigerant
$
.
.
.
.
$
ℎ
$
$
.
.
$
$
.
.
$
$
.
$
Heat Integration
The process underwent heat integration analysis to determine if there were utility savings
possible in the system and to make the system more profitable and efficient. Heat integration
can be incorporated into systems to take advantage of internal heat that is present in one unit
operation and can be used in other unit operations which reduces the need of utilities to
accomplish the required heat duties. Heat integration involves first, identifying potential
streams and unit operations that could be used in a heat exchanger network. Second, the
minimum utility usage was identified from the chosen streams in the first step and the total
152
amount of savings is identified. Finally, the heat exchanger network is designed along with the
duties of each heat exchanger and the improvement over the base case is reported.
Stage 1: Identifying Heat Integration Potential
Beginning in the biomass preprocessing section, the rotary dryer is not included due to fact that
it is a large volume of air at a low temperature and therefore would require a very large heat
exchanger to extract the small amount of usable heat. Also, the air in that temperature range is
needed to dry the corn stover and is not usable. The other units in preprocessing are run
adiabatically and the power requirement does not generate heat and solids are not an efficient
heat transfer medium.
In the gasification of methanol section of the plant, the solar reactor is not included in heat
integration due to fact that the temperature is much higher than elsewhere and is not viable to
exchange heat with any other streams. Also, all pumps and compressors were removed as they
do not have an appreciable heat duty and the power goes into operating the pump with a high
thermal efficiency. The quench tank is included in heat integration as there is a high heat duty
as well as a large temperature range that makes it suitable for heat integration with multiple
streams that require heating. The quench water is not included as the water is the only
available coolant source suitable for the quench tank operation and is treated as a utility. The
amine separator, NH3-“EP HX, is i luded as it has a la ge heat dut a d its te pe atu e a ge
is large and other streams are able to exchange heat with it. The heater between the amine
system and zinc oxide reactor, HEAT1, is included in heat integration as it requires a large duty,
and has a large temperature range that matches with other streams. Also, it is one of the few
unit operations that requires heating instead of cooling utilities and is open to heat integration.
The heat from the methanol reactor, REACT, is included in heat integration as it provides a high
heat duty in a usable temperature range to exchange heat. The cooler between the methanol
reactor and the vapor-liquid separator, VL-SEP HX, is included as well due to having a large heat
duty and is able to exchange heat.
In the amine system, the pumps are not included as their power requirement is to power the
pumps and there is a high level of thermal efficiency. The internal heat exchanger, Lean/Rich
153
Exchanger, is not included because there are no utilities used and is already integrated with the
process. The kettle reboiler for the regenerator column, Reboiler, is included as it is in a usable
temperature range and is one of the few unit operations that requires heating. The
ege e ato s o de se , Co de se , is ot i luded as it uses a
ie t ai i la ge olu es a d
is therefore not in a usable temperature range. The recycle stream cooler, Cooler, is included
though it does have a relatively small heat duty but it has a usable temperature range.
In the methanol purification system, the reboiler for the T-100 column is included as it is in the
temperature range needed and is one of the few streams that requires heating instead of
cooling in our system. The condenser for T-100 is also included as it has a very high heat duty
and is in a usable temperature range to exchange heat. Table 100 below is a summary of which
streams are included in heat integration.
Table 100: List of streams included in Heat Integration Analysis
Stream Name Identifier
Products_To_C-Prod
REACT_Heat
VL-SEP_Heat
NH3-SEP Heat
MeOH Condenser
AG Cooler
CLN-Gas_To_TO-ZN-R
MeOH Reboiler
AG Reboiler
Tin (°C)
1450
279.4
210.0
121.8
80.0
55.4
20.0
129.2
118.3
Tout (°C)
120.0
210.0
50.0
20.0
80.0
43.3
210.0
129.8
120.1
Heat Duty (Btu/hr)
9.48E7
6.26E7
4.60E7
9.16E6
3.74E7
1.93E5
1.30E7
4.09E7
6.00E6
Stage 2: Identifying Minimum Utility Usage and Current Utility Usage
To determine how effective the heat integration will be, a base case for utility usage was
compiled from the simulation. This was found by summing the total heat and cooling duties
performed by the heating and cooling utilities, respectively.
Using the streams in Table 100, an analysis of the minimum utility usages was carried out to
determine how effective heat integration could be in diminishing the use of utilities. The
minimum utility usage is determined using the composite curves in Aspen Energy Analyzer and
this is shown below in Figure 35. The minimum utility rates are found by taking the difference
between the enthalpy of the cold composite curve in blue and hot composite curve in red at
the highest and lowest temperatures, respectively.
154
Figure 35: Aspen Energy Analyzer composite curves for identifying minimum utility usages.
The base case utility usage, and the minimum utility usage calculated from Aspen Energy
Analyzer in Table 101 below. This table also shows the percentage of the base case if the
minimum utility usage was met.
Table 101: Base case utility usage
Utility Type:
Hot Utilities, Base Case
Cold Utilities, Base Case
Hot Utilities, Minimum
Cold Utilities, Minimum
% Savings Possible, Hot
% Savings Possible, Cold
Heat Duty
.
.
.
.
.
%
%
/ℎ
/ℎ
/ℎ
/ℎ
From this, there is no potential to save on hot utility usage but that there is a significant
amount of cooling utilities that could be saved.
155
Stage 3: Creating a Heat Exchanger Network (HEN)
Now that streams have been identified as possibilities for heat incorporatio a d it s ee see
that there are significant opportunities for reducing utility usage, a heat exchanger network is
constructed using Aspen Energy Analyzer. This network will contain a number of new heat
exchangers that will allow the plant to use heat already in the system instead of relying on
constant streams of utilities. Shown below in Table 102 is the heat exchanger network that was
found.
Table 102: Heat Exchanger Network with Associated Heat Duties
Heat Exchanger
E-139
E-140
E-142
E-143
E-144
E-145
Duty (Btu/hr)
.
.
.
.
.
.
Heat Exchanger
E-146
E-147
E-148
E-149
E-150
E-151
Duty (Btu/hr)
.
.
.
.
.
.
This design yields six process heat exchangers and six cold utility heat exchangers with zero hot
utility heat exchangers necessary. Shown below in Table 13 is a summary of the minimum
number of heat exchangers as well as any heat loops in the base case as well as the designed
HEN shown above. Also, the utility percentage of the target is shown. Equation (107) is used to
solve for the minimum number of heat exchangers. The heat exchanger network can be found
in the attached Aspen Energy Analyzer file. A detailed image of the heat exchanger network can
be found in Appendix III-E.
, �
=
+
−
(107)
Table 103: Base Case versus Computed Heat Exchanger Network Performance
HEN Design
Base case
Computed
Actual #
of HX
Minimum #
of HX
Cooling utility usage
(% of target)
. %
%
Heating utility usage
(% of target)
%
%
Heat
loops
156
X. Estimation of Capital Investment and Total Product Cost
Capital Investment
Capital investment refers to the money spent on physical goods and services required to build
the Solar-Thermal Biomass Gasification fa ilit f o
the g ou d up . The total apital
investment estimation requires that each piece of equipment throughout the facility be subject
to a rigorous costing procedure. This procedure results in a dollar amount placed on every piece
of process equipment. Due to the large scale of the facility and high costs of reactors and
pressure vessels, small items of process equipment such as valves and low-capacity pumps
were neglected in this analysis.
Equipment Cost Summary
Methods
The costing of each piece of process equipment throughout the biomass gasification facility was
accomplished primarily by following costing procedures detailed in Product and Process Design
Principles. If additional costing methods were used for specific pieces of equipment, the source
for these methods is given below. Unless otherwise specified, purchase costs do not factor in
the produce price index PPI for capital investments. The PPI for this project is included in later
calculations in the Estimation of Capital Investment and Total Product Cost section. Like in the
Equipment Design section, the Equipment Cost Summary begins with the costing procedure
followed for general equipment present throughout the facility. In-depth costing calculations
may be found in Appendix II-M through Appendix II-X.
Generalized Equipment Costing
Pumps
Costing pumps was divided into two separate costing exercises, one for both the pump itself as
well as the pump motor. The physical pumps were costed by first determining the size factor
( ). The size factor is calculated by the following equation, which takes the total volumetric
flowrate ( ̇ ) in gpm and total pressure head (�
) in feet as inputs:
= ̇∙ �
/
(108)
For centrifugal pumps such as those designed in this biomass gasification process, the sizing
factor is then used in the following base cost (
) calculation:
157
=
[ .
[ln
− .
Following this calculation, the pump-type factor (
]+ .
[ln
]]
) and material factor (
(109)
) are determined
from the Product and Process Design Principles textbook, using tables 22.20 and 22.21
respectively. The pump-type factor is chosen based on total flow rate and pump head range,
while the material factor is chosen based on the pressure and substance passing through the
pump. The purchase cost (
) is then calculated using these parameters in the following way:
=
(110)
The purchase cost calculated above is the purchase cost of the pump itself; the electric motor
powering it must be costed separately. The first step required in this costing procedure is to
multiply the pressure change achieved by the pump by the volumetric flowrate. Pressure units
of Pa and flowrate units of m3/s return a power requirement in watts (W), which can readily be
converted to hp. This value is referred to as the pump brake horsepower (
calculations. The fractional efficiency of the pump (
) in the following
) is calculated using the volumetric
flowrate required by the pump:
=− .
ln ̇ − .
+ .
The fractional efficiency of the electric motor (
= .
+ .
ln ̇
(111)
) is calculated by the following equation:
ln
− .
ln
(112)
While both the fractional efficiencies of the pump and the motor are required to fully specify
these pieces of equipment, for the purposes of costing, only the pump brake horsepower and
motor efficiency are required to obtain a purchase cost:
=
(113)
A detailed description of these calculations can be found in Appendix II-P: Pump Costing.
Heat Exchangers
The majority of heat exchangers throughout the biomass gasification facility were modeled as
shell-and-tube heat exchangers. The costing procedure for these heat exchangers begins with
first determining the shell-and-tube heat exchanger subtype required for a given application.
158
Example heat exchanger types include floating head, fixed head and kettle vaporizers. Kettle
reboilers are more expensive than similarly sized floating head heat exchangers; fixed head
heat exchangers are the most cost-effective. While kettle reboilers are explicitly specified
throughout the process, choosing between fixed or floating head exchangers depends on the
process pressures involved; floating head exchangers are chosen for high pressure applications,
or applications in which condensation or vaporization occurs (Seider, Seader, Lewin, &
Widagdo, 2009).
Once the appropriate heat exchanger type has been determined, the heat exchanger area
(
) in ft2 is used to calculate the base cost of each unit. The base cost formula varies between
heat exchanger types. For the purposes of this example, the equation used to determine
for
a floating head shell-and-tube heat exchanger is shown below.
= ex�{
.
− .
= ex�{
.
− .
Similarly, the equation used to determine
The pressure factor (
[ln
]+ .
[ln
]}
(114)
[ln
]+ .
[ln
]}
(115)
for a kettle reboiler is shown below.
) takes in the shell-side pressure of the heat exchanger being designed in
the following equation, where pressure is input in units of psig:
= .
+ .
(
)+ .
(
)
(116)
The final considerations to be taken into account when determining the purchase cost of a
shell-and-tube heat exchanger are the material factor (
) and length factor ( ). The length
factor can be determined from the tube length using the unnumbered table on page 571 of
Product and Process Design Principles. The material factor takes in constants a and b which are
dependent on shell and tube construction materials respectively. These constants can be
determined from Table 22.25 in Product and Process Design Principles. These material
constants are then placed in the following equation to obtain (
=
+(
)
):
(117)
159
) has units of ft2. The purchase cost of the heat
In this equation, the heat exchanger area (
exchanger is then finally calculated by the following formula:
=
(118)
Detailed calculations for each heat exchanger costed throughout the course of this facility can
be found in Appendix II-V: Heat Exchanger Costing.
Pressure Vessels and Towers
The multiple pressure vessels and towers in the methanol production plant were costed
according to the procedure presented in Product and Process Design Principles (Seider, Seader,
Lewin, & Widagdo, 2009). The critical metrics used in this method were the weight
vessel in lbs., the internal diameter of the vessel
�
of the
in ft, and the seam-to-seam length
of
vertical vessels in ft. This metric was computed during the design of each pressure vessel. The
f.o.b. purchase cost
where
of a pressure vessel was computed by Equation (119) below.
=
is the material factor,
peripherals,
+
(119)
is the purchase cost of the empty vessel and some
is the additional cost of platforms and ladders, and
500. For horizontal vessels where ,
<
.
= ex�{ .
For vertical vessels where ,
For towers where .
= ex�{ .
<
= ex�{ .
<
< ,
<
,
is at a CE cost index of
, Equation (120) was used to find
− .
[ln
]+ .
[ln
] }
(120)
− .
[ln
]+ .
[ln
] }
(121)
[ln
]+ .
[ln
] }
< ,
,
,
, Equation (121) was used to find
, Equation (122) was used to find
− .
.
.
(122)
A few of the pressure vessels designed with this method had weights lower than the range
provided in Product and Process Design Principles – Equations (120) through (122) were
extrapolated out to these lower values of
<
�
<
.
was calculated for horizontal vessels where
using Equation (123):
160
= ,
<
was calculated for vertical vessels where
Equation (124):
=
.
<
was calculated for towers where
(125):
The value for
material factors
=
.
�
�
�
.
(123)
<
<
using
(124)
<
and
<
using Equation
.
.
�
<
.
.
�
<
and
(125)
was set to zero for vessels that fell below these size requirements. The
for the materials utilized in this plant are given for pressure vessels in Table
104 below.
Table 104: Material factors for pressure vessels and towers
Material of construction
�
Carbon steel
1.0
Stainless steel 316
2.1
Monel-400
3.6
For towers equipped with trays or plates, the following method from Product and Process
Design Principles was utilized to determine the f.o.b. purchase cost of trays. The installed cost
for tower trays
where
, including downcomers, was calculated with Equation (126) below:
=
is the number of trays,
(126)
is the number of trays factor,
is the material of construction factor, and
Otherwise,
is the tray type factor,
is the base cost. If
was less than one and was computed by Equation (127):
=
.
.
�
Since all towers designed in the proposed process used sieve trays,
was a function of column diameter
�
≥
,
= .
(127)
= . The material factor
in feet; these functions are given in Table 105 below.
161
Table 105: Material factors for trays
Material of construction
��
Carbon steel
1.0
.
Stainless steel 316
.
Monel
The base cost
+ .
+ .
was determined using Equation (128) below:
=
ex� .
�
�
(128)
�
The costs of mist extractors, inlet diverters, and other small internals for vapor-liquid
separators were accounted for in incidental costs and therefore neglected in the following
calculations.
Cyclones
A single element cyclone unit can be costed according to Equation (129), given from Peters and
⁄ (Peters & Timmerhaus, 2001). This
Timmerhaus, where x is the volumetric gas flow rate in
value would then be multiplied by the number of cyclones and material factor, if necessary.
= .
.
∗
Biomass Pre-Processing
P-1 / Sr-101 Shredding
∗
+ .
∗
+ .
(129)
A direct quote for the rotary knife cutter from S. Howes Inc. was not obtained and therefore
Equation (130) was used (Peters & Timmerhaus, 2001). This equation incorporates the motor,
drive, and guard. The size factor, x, is the capacity, in
=−
.
∗
+
.
.
∗
+
.
(130)
The price was determined to be $16,563 for a single unit. It was determined in the design phase
in section VIII that four rotary knife cutters will be used. The total cost of 4 units is $66,252.
P-2 / RDR-101 Rotary Drying
The rotary drum dryer was costed using Equation (132) below (Peters & Timmerhaus, 2001).
This costing metric assumes that the dryer is using hot air, and the size factor, x, is the
peripheral surface area in
.
162
The design outlined in section VIII has 4 rotary dryers in series. Each of these has a length, L, of
and diameter, D, of
. Using Equation (131) below, the peripheral surface area, Ap,
was calculated for a single unit to be
=
.
(Peters & Timmerhaus, 2001).
=
.
(131)
∗
.
(132)
The purchase price of a single rotary dryer is $152,444. There are 4 units, and the total cost for
this unit operation is $609,776.
P-3 / GR-101 Grinding
A direct quote for the Bliss Industries Hammermill EMF was not obtained and therefore
Equation (133) below was used to determine the purchase price (Seider, Seader, Lewin, &
Widagdo, 2009). This equation assumes the motor and drive is included and the size factor, x, is
the feed rate in
ℎ
. From the simulation, a mass flowrate of 10.22
=
.
∗
.
ℎ
was used.
(133)
Only one unit is required and the purchase price for this unit operation was calculated to be
$67,882.
P-4 / HP-101 Hopper
A lock-hopper has many parts that must be priced independently. First, there is the feeder to
the top hopper, which is designed as a screw-type feeder. The purchase price is determined
from Equation (134) below (Seider, Seader, Lewin, & Widagdo, 2009). This uses a size factor, x,
of the volumetric flowrate in
ℎ
.
,
=
.
∗
.
From the design section, a volumetric flowrate was calculated to be 2,455.3
associated with this is $5,451.
(134)
ℎ
. The cost
163
The top and bottom hoppers are then priced using Equation (135) below (Seider, Seader, Lewin,
& Widagdo, 2009). This assumes carbon steel construction, and the size factor, x, is the volume
in
.
=
,ℎ
.
.
∗
(135)
The volume for the top and bottom hopper was determined from the design in section VIII to
be 818.594
and 40.97
, respectively. The cost for the top and bottom hopper is $14,005
and $3,533, respectively.
The purchasing price of the lock hopper feeding system was determined using the specific price
of a lock-hopper feeder of 88
∗
ℎ
where FIM is Finn Marks and $1 is approximately
equal to 4.00 FIM (Aimo Rautalin, 1992). Using the above volumetric flowrate, the cost of the
feeder system was determined to be $1,529.
The total price for the hopper system is the sum of the above values, which yields a total cost of
the unit operation of $24,518.
Solids Conveying
The solids must be transported between the various unit operations in the biomass preprocessing section. For the rotary knife cutter and rotary dryer steps, belt conveyers are
adequate as the pressure is atmospheric and the particles are large. This is calculated from
Equation (136) below (Peters & Timmerhaus, 2001). This assumes a belt width of 0.75 meters,
and 24 meters was assumed an adequate length.
,
= .
∗
+
.
∗
+
The cost of the belt conveyers was calculated to be $93,048.
.
(136)
The remaining system is designed to use pneumatic conveying to accommodate the micronsized particles and the pressure of 35.0 bar. The purchase price is found using Equation (137)
below where the size factor, x, is the distance conveyed in meters (Peters & Timmerhaus,
2001). The pneumatic conveyance is used to transport the material from the hammermill to the
164
lock-hopper and then from the lock-hopper to the top of the heliostat solar reactor. This unit
operation is 180 meters tall and so it was assumed that 200 meters is adequate.
,
�
=
.
∗
.
(137)
The total cost of the pneumatic conveyance is calculated to be $508,856. The total cost of the
solids conveyance is then calculated to be $601,904.
Summary
The biomass pre-processing capital cost summary is shown below in Table 106.
Table 106: Purchase Price Summary for Biomass Pre-Processing Section
Type of Equipment
P-1 / SR-101 Shredding
P-2 / RDR-101 Rotary Drying
P-3 / GR-101 Grinding
P-4 / HP-101 Hopper
Solids Conveyance
Total for Biomass Pre-Processing:
Purchase Price ($)
66,252
609,776
68,882
24,518
601,998
1,371,426
Solar Field
The cost of the heliostats represents the bulk of the total solar costs due to the type of
materials and large quantity. The heliostat cost is calculated by using the total area of the
heliostats required for all three fields and an estimated average cost for mirrors of $126/m 2.
The CPC is costed using the CPC surface area and is priced at $1260/m2. The tower is costed
using the tower height and Equation (21), referenced in the Project Premises section. Table 107
gives a summary of the solar field costs. Specific costing calculations are located in the
Appendix X.
Table 107: Summary of solar field costs
Type of Equipment
Heliostats
Compound Parabolic Troughs
Solar Tower
Purchase Price ($)
29,236,284
454,507
7,210,636
165
Biomass Gasification
Solar Reactor
In costing the solar reactor, the pressure vessel surrounding the SiC tubes previously described
was costed as a vertical pressure vessel composed of 316 stainless steel using the method
described in detail in Pressure Vessels and Towers. The first SiC tube was priced at $1 million
and each additional tube was priced at $300,000 for a total of $6.1 million.
I dete
i i g the ost of the alu i a li i g, a thi k ess of
a ou d the i side su fa e a ea of
the pressure vessel excluding the three apertures defined a volume of alumina required. Kreith
and Boehm suggested the installed cost of alumina was $2489/m3 in 1981 (Kreith & Boehm,
1985). By adjusting for inflation, the installed price of alumina was determined to be
$6,491.49/m3. This led to a total cost estimate of $173,295.02 for the alumina coating along the
reactor wall. The total tally for the solar reactor was then estimated to be $6,452,049.75.
Zinc Oxide Reactor
The zinc oxide reactor was costed as a single horizontal pressure vessel composed of monel400, as described in the Pressure Vessels and Towers section. The estimated total purchase cost
of the vessel was $411,052.26.
Methanol Reactor
The methanol reactor had a volume of 1094 m3 and the formula used for costing vertical
pressure vessel could not accommodate more than 13,854 ft3 (392.3 m3). In determining an
estimated purchase cost for the reactor, the volume was split into three equal parts and three
equivalent reactors of 12,882 ft3 were priced. The total purchase price for all three reactors was
then determined to be $13,858,564.54.
Spray Quench Tank
The spray quench tank used to cool the product stream exiting the solar reactor was modeled
based on a selective catalytic reduction reductant contact vessel currently in operation at Xcel
E e g s Pa
ee “tatio . Be ause si ila
ate ials e e used i
oth the o igi al “C‘ essel
and the spray quench tank modeled in the current process, a price quote was obtained from
Babcock and Wilcox (B&W), the company responsible for the design of the original SCR vessel.
166
The price quote obtained for this piece of equipment, including operating costs, was slated at
$2.5M when the project was commissioned in 2010.
Heat Exchangers
The procedure followed when costing heat exchangers for the Solar-Thermal Biomass
Gasification facility is explained in detail in the General Equipment Costing section. A summary
of some heat exchanger design parameters used in costing as well as the final purchase price of
the three shell-and-tube heat exchangers used in the biomass gasification subsystem are
provided in the table below.
Table 108. Equipment specification and costing summary of biomass gasification subsystem heat exchangers.
Equipment identifier
Heat transfer
2
area (ft )
Shell-side
Tube
Purchase
pressure (psig)
length (ft)
cost
HEAT-1
1619
450
16
$118,888.83
NH3-SEP HX
791
0
16
$91,306.76
VL-SEP HX
4041
507.6
12
$87,895.21
Compressor
The compressor responsible for elevating the pressure of the process stream prior to entering
the methanol reactor was costed based on power consumption and a material factor. The
following equation was used to determine the base cost (CB), which depends only on power
consumption of the unit. This power consumption was obtained from the Aspen PLUS
simulation, and has units of horsepower (hp) in the following formula:
=
[ .
+ . ln
]
(138)
The base cost is then multiplied by the material factor, FM, to obtain the final purchase cost. The
material factor chosen for this compressor corresponds to stainless steel, which was chosen as
the material of construction due to the high temperatures and pressures of the process stream.
=
(139)
Both the material factor for the compressor and the equations used to obtain the purchase cost
were found in Product and Process Design Principles (Seider, Seader, Lewin, & Widagdo, 2009).
167
The metrics used for determining cost as well as the final purchase cost of the compressor are
given in the following table.
Table 109. Equipment specification and costing summary of biomass gasification subsystem compressor.
Equipment Identifier
Base Cost
Material Factor
Total Purchase Cost
COMPRESS
$2,440,226.00
2.5 (stainless steel)
$6,100,564.99
Pumps
The procedure used for costing pumps throughout the Solar-Thermal Biomass Gasification
facility was elucidated in detail in the General Equipment Costing section. A summary of pump
and motor costs for each pump present in the biomass gasification subsystem is provided in the
following table.
Table 110. Equipment specification and costing summary of biomass gasification subsystem pumps.
Equipment
Pump Material
Pump Cost
Identifier
PUMP-1
Motor
Motor Cost
Total Purchase
Enclosure Type
Ductile iron
$6,766.58
Open, drip-
Cost
$1,756.50
$8,523.08
$3,838.35
$10,383.58
$2,403.78
$9,096.99
proof enclosure
PUMP-2
Ductile iron
$6,545.23
Open, dripproof enclosure
PUMP-3
Ductile iron
$6,693.21
Open, dripproof enclosure
Vapor-Liquid Separators
Two separators were costed in the biomass gasification section of the proposed plant. Each of
these pieces of equipment was costed according to the procedure outlined in the general
Pressure Vessels and Towers costing section (pg. 160). The results of these calculations are
summarized in Table 111 below. More detailed calculations are given in Appendix II-C.
Table 111: Costing results for pressure vessels in the biomass gasification section
Equipment
identifier in PFD
NH3-SEP
VL-SEP
Type of vessel
Horizontal vessel
Weight of
vessel (lbs.)
29515
Platforms and
ladders?
No
Total purchase
cost
$41,303.66
Horizontal vessel
5394
Yes
$56,381.33
168
Cyclone
Both cyclones in the gasification process were costed using Equation (129), referenced in the
Generalized Equipment Costing section and given by Peters and Timmerhaus. The first cyclone
in the process has non-negligible amounts of HCl and H2S, and thus the base cost is multiplied
by a material factor of 2.7 for Monel. The base cost for 4 cyclones is $12,746.32, and with the
Monel material factor included they total to $34, 415.07.
ZN-SPLIT cyclone does not have corrosive compounds entering the system, and thus can be
costed with the base price using carbon steel. For 4 cyclones, this cost was calculated at
$13, 638.36.
Amine Scrubbing
Pressure Vessels and Towers
Four pressure vessels and towers were costed for the amine scrubbing section of the methanol
production plant. Each of these pieces of equipment was costed according to the procedure
outlined in the general Pressure Vessels and Towers costing section (pg. 160). The results of
these calculations are summarized in Table 112 below. More detailed calculations are given in
Appendix II-G: Amine Scrubbing Design.
Table 112: Costing results for pressure vessels and towers in the amine scrubbing section
Equipment identifier in
PFD
Type of vessel
Platforms and
ladders?
Number of
trays
Tower
Weight
of vessel
(lbs.)
29515
Yes
20
Total purchase
cost (tower +
trays)
$273,922.70
Absorber
Separator
Vertical vessel
537
No
-
$18,221.16
Regenerator
Tower
1834
Yes
10
$59,981.65
Regenerator Reflux Drum
Vertical vessel
566
No
-
$18,677.71
Heat Exchangers
Four heat exchangers in the amine scrubbing section were costed. The lean/rich heat exchanger
and the regenerator reboiler were costed as shell-and-tube heat exchangers as done in the
Heat Exchanger Costing section (pg. 158). The lean/rich heat exchanger was chosen to be a
floating head heat exchanger because of the large (>50°C) temperature difference between the
hot and cold streams. The regenerator reboiler was costed as a kettle reboiler because these
169
pieces of equipment are more often used as reboilers in columns than traditional shell-andtube heat exchangers. Important costing metrics, such as the heat transfer area and shell-side
pressure, and the results of these costing calculations are given in Table 113 below. Material
factor constants from Equation (117) for a stainless steel shell and stainless steel tubes are
. and
= .
.
=
Table 113: Costing results for Lean/Rich Exchanger and the reboiler in Regenerator
Equipment identifier
Heat transfer
Shell-side
Tube
Purchase
area (ft2)
pressure (psig)
length (ft)
cost
Lean/Rich Exchanger
311
11.3
12
$95,548.98
Regenerator (reboiler)
384
64.7
12
$115,782.70
The double-pipe heat exchanger was not included in capital cost because the ultimate purchase
cost of the unit was negligibly small. According to Plant Design and Economics for Chemical
Engineers by Peters and Timmerhaus, a stainless steel tube and carbon steel shell double-pipe
heat exchanger for pressures less than 4135 kPa (41.35 bar) can be costed according to
Equation (140) (Peters & Timmerhaus, 2001):
where
.
=
.
.
(140)
is the heat transfer area in m2. With a heat transfer area of
,
=$ ,
.
=
.
=
, which was determined to be negligibly small compared to the cost
of the rest of the equipment.
The air-cooled heat exchanger was costed according to Plant Design and Economics for
Chemical Engineers using its outer surface area
= .
With a surface area of
$
,
.
[ln
−
+
=
] −
.
.
.
.
[ln
=
[ln
] +
.
in m2 and Equation (141):
] +
.
. [ln
[ln
] −
]
.
ln
, the purchase cost was found to be
. It should be noted that this cost includes installation and adjustments for
(141)
=
inflation. These figures were considered in the overall economic analysis.
170
Pumps
The booster pump was costed according to the procedure provided in the Pumps costing
section (pg. 157). The pump was chosen to be made out of stainless steel, making its material
factor
= . The horsepower required for the motor was found to be 6.83 hp with an
efficiency of 88.1%. The motor was costed as a totally enclosed, fan-cooled motor to protect
the motor from the process stream. Motor specifications and overall costs are given in Table
114 below.
Table 114: Costing results for Booster Pump, including motor specifications and cost
Equipment
Purchase cost
Motor
identifier
of pump
power (hp)
Booster Pump
$11,597.49
6.83
Motor type
Totally enclosed,
Purchase cost
Total purchase
of motor ($)
cost ($)
$2,015.63
$13,613.11
fan-cooled
Additional Considerations
A surge tank was neither designed nor costed because storage tanks were accounted for in the
economics analysis as an incidental cost.
Product Separation & Post-Processing
Specific costing calculations for the methanol purification process are located in Appendix II-T.
The valve at the entry of the system was neglected because its cost is negligible in comparison
to the rest of the equipment.
Separator V-100
The V-100 separator unit was costed according to the calculated diameter of 0.76 meter using
an affordable material, carbon steel, due to the lack of corrosive components in the system.
The final cost was calculated to be $191,829.43 using an equation from Peters and Timmerhaus
(Peters & Timmerhaus, 2001).
Distillation Column T-100
The T-100 distillation column was costed using the method of Mulet, Corripio, and Evans as
outlined in the Process and Design textbook. This method provides an f.o.b. purchase cost at a
CE index of 500 and includes an allowance for platforms, ladders, nozzles, and manholes. Table
171
115 shows a summary of these calculations, with the final purchase cost account for a CE index
of 560.
Table 115: Summary of costing specifications for distillation column T-100
Equipment identifier in
PFD
Type of vessel
Distillation T-100
Tower
Weight
of vessel
(lbs.)
65176.47
Platforms and
ladders?
Number of
trays
Yes
34
Total purchase
cost (tower +
trays)
$1,411,732.00
The reboiler and condenser were costed similarly; Table 22.32 was used in the Seider textbook.
The heat transfer area in conjunction with Equation (142) provides the final cost:
=
∗
^ .
(142)
The condenser was costed at $271,098 and the reboiler was costed at $168,794 after adjusting
for a CE index of 560.
Operating Costs
Metrics of Plant Operation
The biomass gasification facility designed here is intended to operate for 8000 hours/year, and
produce a yearly supply of 58,300,000 gallons of highly purified methanol. With 15%
contingencies, the total depreciable capital at startup was found to be $300.5M, while total
working capital, including start-up materials inventory returns $7.93M.
Variable Operating Costs
Operating or variable cost estimates factored utility costs required per year. Utilities are
detailed in the Utility Summary and Heat Integration section. The final value obtained for this
metric was $12.85M annually.
Fixed Operating Costs
Fixed costs factor operator labor. Using operator hiring procedures obtained from Product and
Process Design Principles, plant operators were chosen based on plant subunit (Seider, Seader,
Lewin, & Widagdo, 2009). Biomass pre-processing is estimated to require 6 operators, while
amine scrubbing requires 2 operators. Zinc and methanol reactor operation require 4 operators
each, while the distillation columns and methanol purification subsection require an additional
172
2 operators present at the facility at all times. Due to shift work, total hirable operators require
these values to be summed and multiplied by five in order to account for off days and holiday
time. An additional operator is also kept as a full-time employee to tend to the solar field and
any additional required maintenance. Therefore, the total number of operators required for
plant operation was estimated at 111 full-time employees.
The following table summarizes operator duties and selection.
Table 116. Summary of operator division of labor based on plant equipment and subsection.
Plant Subsystem or Equipment
Biomass Pre-Processing
Amine Scrubbing
Zinc Reactor
Methanol Reactor
Methanol Purification
Solar Field
Number of Operators Required
6
2
4
4
2
1
Total Operators Required:
Multiplying Factor (due to shift
work)
5
5
5
5
5
1
111
Fixed costs including operator labor, benefits, maintenance and overhead totals to $104,000
per operator. The total fixed cost for the biomass gasification facility is $17.55M annually.
173
XI. Profitability Analysis
Profitability
Base Case
The base case profitability analysis models the economic feasibility of the solar-thermal
biomass gasification facility with 12.5% i esto s ate of etu
(IRR). The economic
spreadsheet used to conduct this analysis calculates a number of economic metrics, including
return on investment (ROI), pay-back period (PBP), net present value (NPV) and product price
under given conditions. The commodity market value of the end-product high purity methanol
was determined to be $1.05/gal (Methanex, 2015). Part of the profitability analysis will
determine if this market value can be met while still maintaining 12.5% IRR required to fund the
project. The base case used for this profitability analysis does not factor in any carbon credit
that may be issued to industrial facilities that operate with reduced CO2 emissions. The
calculations for the base case without carbon credits are found in Appendix IV-A.
Cash Flow Analysis
Cash flow analysis of the solar-thermal biomass gasification facility monitors the monetary
influx and outflow throughout the 30-year lifespan of the facility. A cash flow analysis chart
visually displaying this data is shown below. This analysis is performed assuming a one-year
construction period, followed by a startup year in which the plant operates at 50% capacity.
When performing the financial calculations, it was assumed that all costs were incurred in the
first (construction) year. The first construction year was set to occur in 2016, with 2015 serving
as the design year.
174
Cash Flow
$100,000
$50,000
2015
2016
2017
2018
2019
2020
2021
2022
2023
2024
2025
2026
2027
2028
2029
2030
2031
2032
2033
2034
2035
2036
2037
2038
2039
2040
2041
2042
2043
2044
2045
2046
2047
$0
Cash Flow ($k)
-$50,000
-$100,000
-$150,000
-$200,000
-$250,000
-$300,000
-$350,000
Figure 36. Cash flow analysis for 30-year operation of Solar-Thermal Biomass Gasification facility, including construction and
startup.
As illustrated in the above figure, the facility experiences negative cash flow for both the
construction year, and the first year of operation in which the plant only operates at 50%
capacity. Following this, the plant achieves a positive cash flow every year until its end-of-life.
175
Cumulative Cash Flow
1200000
1000000
Cumulative Cash Flow -- $k
800000
600000
400000
200000
2015
2016
2017
2018
2019
2020
2021
2022
2023
2024
2025
2026
2027
2028
2029
2030
2031
2032
2033
2034
2035
2036
2037
2038
2039
2040
2041
2042
2043
2044
2045
2046
2047
0
-200000
-400000
Figure 37. Cumulative cash flow analysis for 30-year Solar-Thermal Biomass Gasification facility lifespan
The cumulative cash flow diagram for the Solar-Thermal Biomass Gasification facility illustrates
that follo i g o st u tio i
, the pla t ill
eak e e fi a iall i
. B the fa ilit s
expected end-of-life, cumulative cash flow will reach a maximum of approximately $1.10B.
Pay Back Period
As is evident in Figure 37, expected pay-back period (PBP) for this facility with 12.5% fixed IRR is
approximately 9.2 years. This value indicates that all money invested in facility construction,
maintenance, personnel costs, taxes and land purchase fees will be recouped prior to the end
of the fi st de ade of pla t ope atio . Afte this pe iod i the pla t s life, the e ai i g i o e
generated by the facility is taken as investor and corporate profit.
Return on Investment
Under the operating conditions stipulated for the Solar-Thermal Biomass Gasification facility,
the expected return on investment for individuals and corporations who contributed to the up176
front costs of plant construction is approximately 10.8%. While this ROI value is higher than one
would expect to receive if that capital had instead been invested in the stock market (Hanlon,
2014), investing instead in a carbon-forward methanol production facility would still be
considered a high-risk investment, due to possible market fluctuations in methanol commodity
price and the possibility of unforeseen technical malfunctions that can occur with a leadingedge chemical production facility.
Net Present Value
The net present value (NPV) is the p ese t alue of the ash flo s at the equired rate of
return of [the Solar-Thermal Biomass Gasification fa ilit ] o pa ed to [the] i itial i est e t
(Gallo, 2014). NPV serves as a metric for project fiscal feasibility alongside pay-back period and
IRR; however, NPV is the only method that takes into account the time value of money (Gallo,
2014). Measuring cash flows at the end of each period, the NPV for the base case Biomass
Gasification facility model is $62.381M. A positive NPV value indicates that the value of the
investment grew over the lifespan of the facility; that the plant is worth more at the end of its
life than the sum of the investment capital put in to building it. The higher the NPV value, the
more promising the investment.
Lang’s Factor
The La g s fa to is a
fa to s that
et i that [esti ates] the apital ost of a he i al pla t usi g o e all
ultipl esti ates of the deli e ed ost of the
(Seider, Seader, Lewin, & Widagdo, 2009). This
ethod
ajo ite s of p o ess e uip e t
ultiplies La g s fa to s
the su
of
purchase costs to obtain total permanent investment (also called fixed capital investment), C TPI,
and total capital investment, CTCI. The total capital investment includes an estimate of working
apital at
% of the total apital i est e t. The La g s Fa to go e i g e uatio s a e gi e
below.
= .
���
∑
= .
� �
∑
�
�
�
�
�
�
�
(143)
(144)
�
177
In this economic analysis, CTCI and CTPI a e k o
, so the La g s fa to s a
e a k-solved for.
Using equations (143) and (144 , the La g s fa to s fo the “ola -Thermal Biomass Gasification
facility were determined to be the following:
���
� �
Product Price
= .
= .
Following thorough data input detailing process variables and fiscal plant requirements,
calculations solved for the required selling price of high purity end-product methanol. In order
to achieve 12.5% fixed IRR for the life of the plant, the end product would need to be sold for
$1.69/gal. This price is not competitive with the current market price of methanol, which is
$1.05/gal (Methanex, 2015).
IRR Variation and Resulting Economic Feasibility Metrics
A final comparison of selling price, ROI, PBP and NPV was conducted in which IRR was varied
with 0%, 5%, 12.5% and 20% IRR, and each resulting metric was measured. These values are
presented in the following table.
Table 117. Variation of IRR and resulting selling price, ROI, PBP and NPV values for Solar-Thermal Biomass Gasification base case
with no carbon credit.
IRR
Selling Price
ROI
PBP
NPV (k)
0.0%
$
0.92
2.5%
40.7
$(170,398.45)
5.0%
$
1.15
4.9%
20.3
$(102,277.32)
12.5%
$
1.69
10.8%
9.2
$ 62,384.58
20.0%
$
2.42
18.5%
5.4
$ 283,391.29
In this IRR variation analysis, the target selling price is only achieved with IRR is reduced
between 0% and 5%. This IRR value is unacceptable for most investors, as the capital required
to fund the project would return more money if invested elsewhere (Hanlon, 2014).
Because the base case with 12.5% fixed IRR could not achieve an end-product selling price
competitive with current market commodity prices, 12.5% IRR is not a feasible fixed metric in
order for the Solar-Thermal Biomass Gasification facility to render the project economically
viable. However, this initial base case does not factor in the possibility of a carbon credit issued
178
to industrial facilities that operate with minimal CO2 emissions. For this reason, an additional
economic analysis was conducted in which a carbon credit was calculated that would both
maintain the required 12.5% fixed IRR and reduce the end-product selling price to commodity
market value.
Modified Base Case: Life Cycle Analysis with Carbon Credit
In order to modify the original economic base case examined above, the option for a carbon
credit was included in subsequent analyses. Here, after the initial product price is returned by
economic analyzing software, an additional credit that returns a dollar amount per pound of
CO2 avoided is added. This dollar amount is adjusted until the product selling price meets
current market value. Detailed calculations for the base case with carbon credit included are in
Appendix IV-B.
After performing the economic analysis, it was determined that a CO2 credit of $0.21/lb is
required in order for the Solar-Thermal Biomass Gasification facility to meet market value for
end-product price.
Cash Flow Analysis
As with the initial base case, a cash flow analysis was performed with the carbon credit case.
The results from this analysis are shown in Figure 38. A cumulative cash flow analysis is shown
in Figure 39.
179
Cash Flow
$100,000
$50,000
Cash Flow ($k)
-$50,000
2015
2016
2017
2018
2019
2020
2021
2022
2023
2024
2025
2026
2027
2028
2029
2030
2031
2032
2033
2034
2035
2036
2037
2038
2039
2040
2041
2042
2043
2044
2045
2046
2047
$0
-$100,000
-$150,000
-$200,000
-$250,000
-$300,000
-$350,000
Figure 38. Cash flow analysis for Solar-Thermal Biomass Gasification facility with carbon credit.
Cumulative Cash Flow
1200000
800000
600000
400000
200000
0
-200000
2015
2016
2017
2018
2019
2020
2021
2022
2023
2024
2025
2026
2027
2028
2029
2030
2031
2032
2033
2034
2035
2036
2037
2038
2039
2040
2041
2042
2043
2044
2045
2046
2047
Cumulative Cash Flow -- $k
1000000
-400000
Figure 39. Cumulative cash flow analysis for Solar-Thermal Biomass Gasification Facility with carbon credit.
In both annual and cumulative cash flow analysis diagrams, fiscal trends for the carbon credit
case mimic those in the base case, with only minor differences. Both cases see a financial breakeven occur in 2024, with approximately $1.1B in peak total cumulative cash flow. However,
180
while the no-carbon credit case sees a fiscal deficit in 2017 (Figure 36), the carbon credit case
sees this as the first year of profit (Figure 38).
Pay Back Period and Return on Investment
Despite small changes in the annual and cumulative cash flows, no change in ROI and PBP were
observed via the addition of a carbon credit. This makes sense, since the carbon credit
functions to subsidize the price of the end product, not the capital investment itself.
Net Present Value
The NPV in the carbon credit case was calculated to be $61.629M, using cash flows at the end
of each period. This value is slightly lower than the NPV for the initial base case, however the
difference observed is relatively small.
Product Price
The primary purpose of conducting the second economic analysis with carbon credit was to
reduce the selling price of methanol end-product to meet commodity market value while still
maintaining a 12.5% fixed IRR. This metric was achieved with $0.21/lb carbon credit.
IRR Variation and Resulting Economic Feasibility Metrics
As with the base case, the carbon-credit case was subject to the same IRR variation, in with IRR
percentages of 0%, 5%, 12.5% and 20% were used in economic calculations of selling price, ROI,
PBP and NPV. These values are presented in the following table.
Table 118. Effects of IRR variation on selling price, ROI, PBP and NPV for carbon-credit economic analysis.
IRR
Selling Price
ROI
PBP
NPV
0.0%
$
0.29
2.4%
41.1
$(167,549.44)
5.0%
$
0.52
4.9%
20.4
$(100,641.29
12.5%
$
1.05
10.8%
9.3
$ 61,467.28
20.0%
$
1.77
18.5%
5.4
$ 279,568.97
Sensitivity Analysis
As part of the economic analysis conducted for the initial base case, a sensitivity analysis was
pe fo
ed. This se siti it a al sis e a i ed the effe ts of pe tu i g the fa ilit s total apital
181
expenditure (TCI), fixed operating cost (FOC), variable cost (VC), price of natural gas, price of
heliostats and price of biomass on the return on investment (ROI) and internal rate of return
(IRR) respectively. The results from this sensitivity analysis are displayed in the following
figures.
0.25
0.2
0.2
0.15
0.15
0.1
0.1
IRR
ROI
0.25
ROI
IRR
0.05
0.05
0
0
Total Capital Expenditure
Figure 40. Sensitivity analysis examining the result of perturbations in Total Capital Expenditure on the investment ROI and IRR
for the economic base case.
The first comparison performed in the sensitivity analysis charts the change in ROI and IRR with
perturbations in total capital expenditure. As seen in Figure 40, both the ROI and IRR decrease
exponentially with increases in TCE. This indicates that for the Solar-Thermal Biomass
Gasifi atio fa ilit p oje t, i esto s ‘OI a d I‘‘ a e
a i ized he TCE is
i i ized. I
order to increase the economic viability of the project, further analysis may be performed to
minimize the capital dollars spent on the project design.
182
0.13
0.15
0.14
0.12
0.13
0.11
0.1
0.11
0.1
IRR
ROI
0.12
ROI
IRR
0.09
0.09
0.08
0.08
0.07
5000
10000
15000
20000
25000
0.07
30000
Fixed Operating Cost
Figure 41.Sensitivity analysis examining the result of perturbations in Fixed Operating Cost on the investment ROI and IRR for the
economic base case.
Sensitivity analysis performed on fixed operating costs (FOC) for the Solar-Thermal Biomass
Gasification facility show a linear decrease in ROI and IRR with increases in FOC. Fixed operating
costs include operator labor, benefits and maintenance. Because operator compensation is a
fixed value per employee, maximizing ROI and IRR can only be accomplished by hiring fewer
operators. Because it is unlikely that hiring fewer operators to run a large, leading-edge facility
is a viable option, raising ROI and IRR by these means may be difficult.
183
0.14
0.16
0.15
0.13
0.14
0.12
0.13
0.12
0.1
0.11
IRR
ROI
0.11
ROI
0.1
0.09
IRR
0.09
0.08
0.08
0.07
0.06
$0.10
0.07
$0.20
$0.30
$0.40
$0.50
$0.60
0.06
$0.70
Variable Cost
Figure 42.Sensitivity analysis examining the result of perturbations in Variable Cost on the investment ROI and IRR for the
economic base case.
Variable cost considerations factor in costs of utilities, chemicals and catalysts as well as ash
credit. For the base case economic analysis, no carbon credit was given for the Solar-Thermal
Biomass Gasification facility; however, such credit would reduce the variable cost associated
with the facility. As expected, Figure 42 illustrates a linear decline in both ROI and IRR with
i
easi g a ia le ost. The efo e,
the pla ts a ia le ost, hi h
a i izi g i esto s etu
is a o plished
a o u i the e e t of the issui g of a a o
lo e i g
edit.
184
0.116
0.134
0.114
0.132
0.13
0.112
0.128
0.11
IRR
ROI
0.126
0.108
0.124
ROI
IRR
0.106
0.122
0.104
0.12
0.102
0.118
0.1
0
0.0005
0.001
0.0015
0.002
0.0025
0.003
0.116
0.0035
Price of Natural Gas ($)
Figure 43. Sensitivity analysis examining the result of perturbations in Natural Gas Price on the investment ROI and IRR for the
economic base case.
The price of natural gas strongly influences both ROI and IRR. As natural gas price increases, ROI
and IRR decrease linearly, as shown in Figure 43. Natural gas is an expensive utility. For the
purposes of this design report, the cost of natural gas was estimated at $2/1000 SCF. For this
reason, it is not surprising that increases in natural gas price result in dramatic downturns in
ROI and IRR for the project. Unfortunately, natural gas prices are subject to market fluctuations,
and possible downturns in IRR and ROI resulting from increased commodity prices are likely
difficult to avoid or mitigate.
185
0.14
0.135
0.135
0.13
0.13
0.125
0.125
0.12
0.12
0.115
0.115
0.11
0.11
0.105
0.105
IRR
ROI
0.14
ROI
IRR
0.1
0.1
0.095
0.095
0.09
0
5000
0.09
10000 15000 20000 25000 30000 35000 40000 45000 50000
Price of Heliostats ($)
Figure 44. Sensitivity analysis examining the result of perturbations in Heliostat Price on the investment ROI and IRR for the
economic base case.
The base cost of heliostats was assumed to be $126/m2. Perturbations in this price result in
e pe ted shifts i ‘OI a d I‘‘; i
eases i heliostat p i e de ease i esto s etu , hile
lowering this base price results in slightly higher ROI and IRR. In light of this trend, attempts to
create a master service agreement (MSA) with a heliostat manufacturer may lower the unit
p i e, the e
i
easi g i esto s etu
o the fa ilit .
186
0.14
0.135
0.135
0.13
0.13
0.125
0.125
0.12
0.12
0.115
0.115
0.11
0.11
0.105
0.105
IRR
ROI
0.14
ROI
IRR
0.1
0.1
0.095
0.095
0.09
0.09
0
10
20
30
40
50
60
70
80
90
100
Price of Biomass ($)
Figure 45. Sensitivity analysis examining the result of perturbations in Biomass Price on the investment ROI and IRR for the
economic base case.
As illustrated in Figure 45, increasing the commodity price of biomass results in linear decreases
in ROI and IRR. Because biomass is the primary reagent purchased in order to fuel the SolarThermal Biomass Gasification facility, it is unsurprising that perturbations in this commodity
p i e i flue e i esto s etu s. Ho e e , i
o pa iso to heliostat p i e Figure 44) or
natural gas price (Figure 43), large changes in the commodity price of biomass result in
comparatively small fluctuations in ROI and IRR. This indicates a promising economic resiliency
of the gasifi atio fa ilit , i that i esto s etu s a e ot st o gl depe de t o s all
a ket
fluctuations in primary reagent cost.
187
XII. Conclusion
While the Solar-Thermal Biomass Gasification facility modeled here is technically promising and
environmentally revolutionary, it currently lacks economic feasibility. The inability for the
process, as currently designed, to produce a high-purity methanol end product with a selling
price comparable to current market commodity prices renders the project unlikely to be fiscally
viable. With the fixed 12.5% IRR and observed 10.8% ROI investors can expect from the project
are attractive economic metrics, sensitivity analyses reveal that these values are strongly
dependent on a number of market factors and commodity prices that exist outside of the scope
of control of plant designers. These fiscal uncertainties detract from the economic
attractiveness of the project as a whole.
Despite these economic pitfalls, preserving the environment and striving to achieve a balance
between ecosystems and industry is a noble goal for any chemical manufacturer. Increasingly,
politicians and government agencies have moved toward encouraging lower CO2 emissions
from the largest producers of this environmental pollutant. If a carbon credit is issued to
facilities that operate in an environmentally-friendly way, this credit might make the current
project modeled here increasingly economically viable. If this becomes the case, the design
proposed here may serve as a template for a greener way to produce high-purity commodity
methanol.
188
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195
XIII. Appendix
Appendix I: Engineering Calculations
Appendix I-A: Approach Calculations
196
197
198
199
Appendix I-B: Material and Energy Balances
Aspen PLUS:
200
201
Biomass Pre-Processing
202
Amine Scrubbing
203
Methanol Purification System
204
Appendix I-C: Split Fraction for AG-CLEAN
205
Appendix I-D: Adiabatic Temperature Rise for Methanol Reactor
206
Appendix II: Design and Costing Spreadsheets
Appendix II-A: Solar Field Design
207
Appendix II-B: Cyclone Design
208
Appendix II-C: Aspen PLUS Separator Design
209
210
Appendix II-D: Pump Design
211
212
Appendix II-E: Compressor Design
213
Appendix II-F: Pre-Processing Design
214
215
Appendix II-G: Amine Scrubbing Design
216
217
218
219
220
221
222
223
Appendix II-H: Methanol Purification System Design
224
225
226
227
Appendix II-I: Quench Tank Design
228
229
Appendix II-J: Heat Exchanger Design
230
231
232
233
234
Appendix II-K: Reactor Heat Exchanger Design
235
Appendix II-L: Reactor Design
Solar Reactor
T
1723.15 K
R
8.314 J/K-mol
k1
E1
2510 bar s
112.6 kJ/mol
K1
0.96884601 bar s
k2
E2
0.0674 bar
-37.3 kJ/mol
K2
k3
E3
K3
ϕ
-1 -1
-1 -1
-1
0.910736668 bar
-1
-1
0.304 bar
-36.6 kJ/mol
3.911887174 bar
-1
4.3
source for methane reforming: http://ac.els-cdn.com/S10039
kSRM1
2.924548469
kmol/(g cat-s-bar)
KpSRM1
1993271.267
--
KCO
5.9385E-07
bar
KH2
1.87791E-11
bar
KCH4
0.009621944
bar
KH2O
3.62809E-08
--
kSRM2
1.171040028 kmol ·bar
0.5
KpSRM2
1004719.329
--
-1
-1
-1
/(g cat-s)
236
kWGS
0.000329966 kmol ·bar
0.5
KpWGS
0.504055492
--
kDRM
KpDRM
/(gcat-s)
2
3.40050673 kmol/(g cat-s-bar )
5.802036715
--
K1
0.5
bar
K2
9.71
bar
K3
26.21
bar
gcat
-1
-1
-1
1 g of Ni-CeO2/MgAl 2O*essentially
under uncatalyzed
4
237
pH2 (bar)
pCO (bar) pH2O (bar) pCO2 (bar) pCH4 (bar) X
k0
dX/dt (s-1) rSRM1
rSRM2
rWGS
rDRM
10.08965
1.881945
12.35677
0.000453
7.702292
0.492084
0.231458
0.232552
0.746087
3.691513
-0.00041
-9.03018
10.09023
10.08965
1.893753
1.905203
12.34436
12.33028
0.003502
0.006404
7.704026
7.704517
0.494409
0.496728
0.231266
0.23107
0.231877
0.231197
0.745375
0.744674
3.684291
3.676622
-0.00041
-0.00042
-8.43038
-7.93508
10.0893
10.08914
1.916567
1.927844
12.31629
12.3024
0.009186
0.011866
7.704892
7.70517
0.49904
0.501345
0.230871
0.23067
0.230512
0.229822
0.743925
0.743138
3.668758
3.660743
-0.00042
-0.00042
-7.51961
-7.16482
10.08914
10.08928
1.939035
1.95014
12.28859
12.27488
0.014459
0.016975
7.705367
7.705492
0.503644
0.505935
0.230467
0.230263
0.22913
0.228435
0.74232
0.741476
3.652609
3.644384
-0.00042
-0.00043
-6.85744
-6.58793
10.08955
10.08992
1.961159
1.972093
12.26126
12.24773
0.019425
0.021816
7.705557
7.705567
0.508219
0.510497
0.230058
0.229853
0.227737
0.227037
0.740609
0.739725
3.636087
3.627737
-0.00043
-0.00043
-6.34922
-6.13597
10.09039
10.09094
1.982942
1.993707
12.23429
12.22095
0.024153
0.026442
7.705529
7.705449
0.512767
0.51503
0.229647
0.229441
0.226335
0.225632
0.738825
0.737911
3.619346
3.610927
-0.00044
-0.00044
-5.94403
-5.77015
10.09157
10.09226
10.09302
10.09383
2.004387
2.014984
2.025497
2.035928
12.20769
12.19452
12.18144
12.16845
0.028687
0.030892
0.033059
0.035193
7.70533
7.705176
7.704992
7.704778
0.517287
0.519536
0.521778
0.524013
0.229235
0.229028
0.228823
0.228617
0.224926
0.22422
0.223512
0.222804
0.736986
0.736052
0.735109
0.73416
3.602489
3.594041
3.585589
3.577139
-0.00044
-0.00045
-0.00045
-0.00045
-5.61172
-5.46664
-5.33316
-5.20985
10.0947
10.09561
10.09656
10.09756
10.09859
10.09966
2.046276
2.056543
2.066727
2.076831
2.086854
2.096797
12.15555
12.14274
12.13002
12.11738
12.10483
12.09237
0.037294
0.039366
0.041411
0.043429
0.045423
0.047393
7.704538
7.704274
7.703988
7.703681
7.703355
7.703012
0.526241
0.528462
0.530676
0.532883
0.535082
0.537275
0.228412
0.228207
0.228002
0.227798
0.227595
0.227392
0.222094
0.221384
0.220673
0.219961
0.219249
0.218536
0.733205
0.732245
0.731281
0.730314
0.729344
0.728373
3.568698
3.560269
3.551857
3.543464
3.535095
3.526751
-0.00046
-0.00046
-0.00046
-0.00047
-0.00047
-0.00047
-5.09551
-4.98914
-4.88986
-4.79693
-4.70973
-4.6277
10.10076
2.10666
12.07999
0.049342
7.702652
0.53946
0.22719
0.217823
0.7274
3.518436
-0.00047
-4.55035
10.10189
2.116443
12.0677
0.05127
7.702277
0.541638
0.226989
0.217109
0.726427
3.510151
-0.00048
-4.47728
10.10304
2.126148
12.05549
0.053179
7.701888
0.54381
0.226789
0.216395
0.725453
3.501899
-0.00048
-4.4081
10.10422
2.135774
12.04337
0.055069
7.701486
0.545973
0.226589
0.215681
0.724479
3.49368
-0.00048
-4.34248
10.10543
2.145323
12.03133
0.05694
7.701071
0.54813
0.22639
0.214967
0.723505
3.485497
-0.00049
-4.28015
10.10665
2.154794
12.01937
0.058795
7.700644
0.55028
0.226191
0.214253
0.722532
3.477351
-0.00049
-4.22083
10.1079
10.82226
10.82358
10.82491
10.82623
10.82755
10.82887
10.83019
10.83151
10.83283
10.83415
10.83546
10.83678
10.8381
10.83941
10.84073
2.164188
3.139238
3.13928
3.139321
3.139362
3.139402
3.139442
3.139482
3.139521
3.13956
3.139598
3.139636
3.139673
3.13971
3.139747
3.139783
12.0075
10.17319
10.17174
10.1703
10.16886
10.16742
10.16598
10.16454
10.16311
10.16167
10.16024
10.15881
10.15738
10.15595
10.15453
10.1531
0.060634
0.559777
0.560515
0.561253
0.56199
0.562726
0.563462
0.564197
0.564932
0.565666
0.566399
0.567132
0.567864
0.568595
0.569326
0.570056
7.700206
7.405669
7.405107
7.404546
7.403985
7.403424
7.402863
7.402303
7.401743
7.401183
7.400624
7.400065
7.399506
7.398947
7.398389
7.397831
0.552422
0.979591
0.979752
0.979912
0.980071
0.980228
0.980385
0.98054
0.980695
0.980848
0.981
0.981151
0.981301
0.981449
0.981597
0.981744
0.225994
0.187379
0.187339
0.187299
0.187258
0.187218
0.187178
0.187138
0.187098
0.187057
0.187017
0.186977
0.186937
0.186897
0.186858
0.186818
0.213539
0.016105
0.015989
0.015875
0.015761
0.015648
0.015536
0.015424
0.015313
0.015203
0.015093
0.014985
0.014876
0.014769
0.014662
0.014556
0.72156
0.498293
0.498038
0.497782
0.497527
0.497272
0.497017
0.496763
0.496509
0.496255
0.496002
0.495749
0.495496
0.495244
0.494991
0.49474
3.469242
2.029749
2.028418
2.02709
2.025764
2.024439
2.023116
2.021795
2.020476
2.019159
2.017843
2.016529
2.015217
2.013907
2.012599
2.011292
-0.00049
-0.00094
-0.00094
-0.00094
-0.00094
-0.00094
-0.00094
-0.00094
-0.00094
-0.00094
-0.00094
-0.00094
-0.00094
-0.00095
-0.00095
-0.00095
-4.16429 *row 30
-1.20124 *row514
-1.20005
-1.19887
-1.19769
-1.19652
-1.19535
-1.19418
-1.19302
-1.19186
-1.19071
-1.18956
-1.18841
-1.18727
-1.18613
-1.18499
10.84204
3.139819
10.15168
0.570786
7.397273
0.981889
0.186778
0.01445
0.494488
2.009988
-0.00095
-1.18386
238
Methanol Reactor
239
Zinc Oxide Reactor
240
Appendix II-M: Solar Field Costing
241
Appendix II-N: Cyclone Costing
242
Appendix II-O: Aspen PLUS Separator Costing
243
Appendix II-P: Pump Costing
244
245
246
Appendix II-Q: Compressor Costing
247
Appendix II-R: Pre-Processing Costing
248
249
Appendix II-S: Amine Scrubbing Costing
250
251
Appendix II-T: Methanol Purification System Costing
252
Appendix II-U: Quench Tank Costing
253
Appendix II-V: Heat Exchanger Costing
254
255
Appendix II-W: Reactor Heat Exchanger Costing
256
Appendix II-X: Reactor Costing
257
Appendix III: Computer Process Modeling and Simulations
Appendix III-A: Biomass Pre-Processing Simulation
258
Appendix III-B: Aspen PLUS Simulation
R- SOLAR
SOLAR QN CH- H2O
BRKD OW N
CALCULATOR
H2O -O UT
PUM P-1
METH ANE
DEC OMP
PRO
DUC
MIX-H
2O TS
PUM P-2
BIO MASS
TO- ZN- R
ACID- GAS
VAPOR
SPRAY- Q
SPLIT- 1
ZN- VAP
PUR GE
SPLIT- 2
TO- SOLAR
TO- CO MP
SEP- VAP
MIX-1
H2O -R CT
MIX-2
COZNMPRESS
SOLID
C- PRO D
ZN- PRO D
ZN- SPLIT
REAC T
ZN- REAC T
HEAT1
CLN -G AS
TO- CYC L
CYC LONE
AG- CLEAN
VL- SEP
NH 3-SEP
RC T-PR OD
H2O -3
ACIDS
PUM P-3
REC YCLE
XTRA- H20
259
Appendix III-C: Amine Scrubbing Simulation
260
Appendix III-D: Methanol Purification Simulation
261
Appendix III-E: Heat Integration
262
Appendix III-F: Reactor Computer Code
Methanol Reactor Polymath
T = 210+273 #K
R1 = 8.314 #J/K-mol
PM = M/(M+CO+CO2+H2+235.1162525)*78.95386 #atm
PH2 = H2/(M+CO+CO2+H2+235.1162525)*78.95386 #atm
PCO2 = CO2/(M+CO+CO2+H2+235.1162525)*78.95386 #atm
PCO = CO/(M+CO+CO2+H2+235.1162525)*78.95386 #atm
PH2O = 146/(M+CO+CO2+H2+235.1162525)*78.95386 #atm
gcat= 1094.027/1000*25/0.0005*25 #kg -- 1094.027 m3/(1000 g/kg)(25 g/.0005 m3)*25 volumes/volume
(more concentrated than in smaller scale)
rootKH2 = 0.499*exp(17197/R1/T)
KH2O = 6.62*10^-11*exp(124119/R1/T)
KH2OoK8K9KH2 = 3453.38
k5aK2K3K4KH2 = 1.07*exp(36696/R1/T)
k1 = 1.22*10^10*exp(-94765/R1/T)
Oneok3star = 10^(-2073/T+2.029)
rRWGS = gcat*k1*PCO2*(1(1/Oneok3star)*(PH2O*1.01325*PCO*1.01325/(PCO2*1.01325)/(PH2*1.01325)))/(1+(KH2OoK8K9KH2)*(
PH2O*1.01325/(1.01325*PH2))+rootKH2*((1.01325*PH2)^0.5)+KH2O*PH2O*1.01325)*(60*60/1000)
#kmol/hr
E = 18360.0 #cal/mol
Ko = 965.96 #mol/(kg cat-hr)
K1 = 15.0019*(10^-3) #unitless
K2 = 1.488*(10^-3) #unitless
K3 = 3.957*(10^-3) #unitless
K4 = 0.03677*(10^-3) #unitless
R = 1.9872035 #cal/K-mol
COF1 = 3927/T-7.971*log(T)+0.002499*T-2.953*(10^-7)*T^2+10.2 #Vapor phase equilibrium constant
KA = 10^COF1
HH2 = exp(-11.1158+1438.0219/T+1.9043*ln(T)) #Henry's Law coeff for H2 MPa (kmol solvent/kmol
solute): http://pubs.acs.org/doi/pdf/10.1021/ef00002a012
HCO = exp(88.9926-6417.1251/T-11.6310*ln(T)) #Henry's Law coeff for CO MPa (kmol solvent/kmol
solute): Ko et al
HM = 195*exp(5400*(1/T-1/298.15))/0.10325 #Henry's Law coeff for CH3OH MPa (kmol solvent/kmol
solute): http://www.henrys-law.org/henry-3.0.pdf
#agrees fairly closely with experimental data obtained from Ko et al. for this case
KL = KA*HH2^2*HCO/HM*(9.86923^2) #atm^2
a = 0.818
b = 0.82323
c = 2.0903
d = 2.1598
rMeOH = (gcat*Ko*exp(-E/(R*T))*(PH2^2*PCO-PM/KL)/(K1+K2*PH2^a+K3*PCO^b+K4*PM^c)^d)/1000
#kmol/hr
d(H2) / d(t) = -rMeOH #kmol/hr
263
H2(0) = 2886.79 #kmols/hr
d(CO2) / d(t) = -rMeOH-rRWGS #kmol/hr
CO2(0) = 1.967449 #kmols/hr
d(CO) / d(t) = rRWGS #kmol/hr
CO(0) = 1427.629
d(M) / d(t) = rMeOH #kmol/hr
M(0) = 30.72034 #kmols/hr
t(0)=0
t(f)= 0.02 #hr
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Solar Reactor FORTRAN
BRKDOWN:
C
FACT IS THE FACTOR THAT CONVERTS THE FEED TO A WET BASIS.
FACT = (100 - WATER) / 100
H2O
= WATER
/ 100
ASH
= ULT(1) / 100 * FACT
CARB = ULT(2) / 100 * FACT
H2
= ULT(3) / 100 * FACT
N2
= ULT(4) / 100 * FACT
CL2
= ULT(5) / 100 * FACT
SULF = ULT(6) / 100 * FACT
O2
= ULT(7) / 100 * FACT
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Appendix IV: Economic Analysis Spreadsheets
Appendix IV-A: Base Case without Carbon Credit
266
267
268
269
270
271
272
273
Appendix IV-B: Base Case with Carbon Credit
274
275
276
277
278
279
280
281
Appendix IV-C: Summary of Utilities
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