United States
Environmental Protection
Agency
Municipal Environmental Research
Laboratory
Cincinnati OH 45268
EPA-600 2 80 122
August 1980
Research and Development
6EPA
Pyrolytic Oils
Characterization and
Data Development for
Continuous Processing
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RESEARCH REPORTING SERIES
Research reports of the Office of Research and Development, U.S. Environmental
Protection Agency, have been grouped into nine series. These nine broad cate-
gories were established to facilitate further development and application of en-
vironmental technology. Elimination of traditional grouping was consciously
planned to foster technology transfer and a maximum interface in related fields.
The nine series are:
1. Environmental Health Effects Research
2. Environmental Protection Technology
3. Ecological Research
4. Environmental Monitoring
5. Socioeconomic Environmental Studies
6. Scientific and Technical Assessment Reports (STAR)
7. Interagency Energy-Environment Research and Development
8. "Special" Reports
9. Miscellaneous Reports
This report has been assigned to the ENVIRONMENTAL PROTECTION TECH-
NOLOGY series. This series describes research performed to develop and dem-
onstrate instrumentation, equipment, and methodology to repair or prevent en-
vironmental degradation from point and non-point sources of pollution. This work
provides the new or improved technology required for the control and treatment
of pollution-sources to meet environmental quality standards.
This document is available to the public through the National Technical Informa-
tion Service, Springfield, Virginia 22161.
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EPA-600/2-80-122
August 1980
PYROLYTIC OILS - CHARACTERIZATION AND DATA
DEVELOPMENT FOR CONTINUOUS PROCESSING
by
J. A. Knight, L. W. Elston,
D. R. Hurst, and R. J. Kovac
Engineering Experiment Station
Georgia Institute of Technology
Atlanta, Georgia 30332
Grant Nos. R-804416 and R-806403
Project Officer
Charles J. Rogers
Solid and Hazardous Waste Research Division
Municipal Environmental Research Laboratory
Cincinnati, Ohio 45268
MUNICIPAL ENVIRONMENTAL RESEARCH LABORATORY
OFFICE OF RESEARCH AND DEVELOPMENT
U. S. ENVIRONMENTAL PROTECTION AGENCY
CINCINNATI, OHIO 45268
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DISCLAIMER
This report has been reviewed by the Municipal Environmental Research
Laboratory, U. S. Environmental Protection Agency, arid approved for publica-
tion. Approval does not signify that the contents necessarily reflect the
views and policies of the U. S. Environmental Protection Agency, nor does
mention of trade names or commercial products constitute endorsement or recom-
mendation for use.
ii
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FOREWORD
The U. S. Environmental Protection Agency was created because of
increasing public and government concern about the dangers of pollution to
the health and welfare of the American people. Noxious air, foul water, and
spoiled land are tragic testimonies to the deterioration of our natural
environment. The complexity of that environment and the interplay of its
components require a concentrated and integrated attack on the problem.
Research and development is that necessary first step in problem solu-
tion; it involves defining the problem, measuring its impact, and searching
for solutions. The Municipal Environmental Research Laboratory develops new
and improved technology and systems to prevent, treat, and manage wastewater
and solid and hazardous waste pollutant discharges from municipal and comm-
unity sources, to preserve and treat public drinking water supplies, and to
minimize the adverse economic, social, health, and aesthetic effects of pollu-
tion. This publication is one of the products of that research and provides
a most vital communications link between the researcher and the user community.
This is a report on the characterization of oils obtained by the
pyrblysis of lignocellulosic wastes and the development of processing tech-
niques that would yield fractions suitable for industrial applications.
iii
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ABSTRACT
Pyrolytic oils produced by the pyrolysis of forestry residues in a ver-
tical bed, countercurrent flow reactor (Georgia Tech pyrolysis process) have
been thoroughly characterized. The pyrolytic oils were produced in a 500 Ib
per hour pilot plant and in a 50 ton per day field development facility. The
overall chemical and physical properties have been determined by standard ana-
lytical techniques. The oils are dark brown to black with a burnt, pungent
odor and have a boiling range of about 100°C to approximately 200°C at which
point thermal degradation begins to occur. The heating values of the oils,
which burn cleanly, are approximately two-thirds of petroleum fuel oil heating
values. The oils, which are acidic, exhibit some corrosive characteristics.
The oils are composed of a large number of oxygenated compounds which exhibit
a wide spectrum of chemical functionality. Based on the results of this study,
the pyrolytic oils contained phenolics, polyhydroxy neutral compounds, neutral
compounds of a high degree of aromaticity and volatile acidic compounds.
A number of approaches to separating the oils into fractions, each of
which would contain a predominant chemical species, were investigated on a
batch basis. These approaches employed extraction techniques with water,
organic solvents, aqueous alkaline solutions, and aqueous salt solutions.
Based on the experimental results on a batch basis, two approaches were selec-
ted for continuous extraction experiments at the bench level with both raw oil
and vacuum stripped oil. The results of these continuous extraction experi-
ments show that these approaches are very promising as processing methods for
producing oil fractions which would be useful for industrial chemical applica-
tions. Based on the results of the continuous extraction experiments, a ver-
satile pilot plant was designed for further investigation of pyrolytic oils
which would yield data for scale up of the process for a commercial plant and
produce oil fractions for studies for industrial applications. Preliminary
economic assessments, based on two approaches, indicate that the processing of
pyrolytic oils could be economically viable. The results indicate that, for a
50 percent net return on investment, the selling price for the oil fractions
would have to be in the range of 8.4 to 10.6 cents per pound which is in the
same range as 9 cents per pound for coal tar creosote and well below 54 cents
per pound for cresylic acid, which were quoted market prices in December, 1979.
The preliminary economic assessments are encouraging for processing pyrolytic
oils into fractions suitable for industrial chemical applications.
This report was submitted in fulfillment of Grant Nos. R-804416 and
R-806403 by Georgia Institute of Technology under the sponsorship of the U. S.
Environmental Protection Agency. This report covers the period June 21, 1976
to March 31, 1980, and work was completed as of March 31, 1980.
iv
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CONTENTS
/ Page
Foreword ill
Abstract . .......... iv
Figures vi
Tables -. ix
Acknowledgment xii
1. Introduction 1
2. Summary < 3
3. Recommendations 6
4. Background Information . 7
5. Experimental 11
Phase I ' 11
Phase II 38
Phase III 52
6. Pilot Plant Design 77
7. Design and Economics of Commercial Size Plant 88
8. Discussion 116
References 126
Appendices
A. Material Balance Calculations ... 128
B. Pilot Plant Calculations . 137
C. Commercial Plant Calculations 159
D. Physical Properties 181
Glossary 184
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FIGURES
Number Page
1 Viscosity of condenser oil 16
2 Viscosity of draft fan oil 16
3 Vacuum stripped condenser oil 17
4 Vacuum stripped draft fan oil 18
5 Effect of heating condenser oil at 110°C for
different time periods on viscosity 19
6 Viscosity curves for condenser oil (initial) and
No. 2 and No. 6 fuel oils 20
7 Liquid chromatogram of wood oil. Partisil PAC column with
0-100% solvent gradient of 2-propanol in iso-octane 22
8 Liquid chromatogram of wood oil. Partisil ODS column with
10-100% solvent gradient of acetonitrile in water 22
9 Liquid chromatogram of wood oil. Partisil ODS column
with 10-100% solvent gradient of acetonitrile in
water with 20 minute hold at 40% acetonitrile . 23
10 Liquid chromatogram of wood oil at 210 nm 25
11 Liquid chromatogram of wood oil at 254 nm 25
12 Liquid chromatogram of wood oil at 280 nm 26
»
13 Liquid chromatogram of wood oil at 300 nm 26
14 Liquid chromatogram of wood oil at 360 nm 27
15 Survey liquid chromatogram of raw condenser oil 28
16 Survey liquid chromatogram of draft fan oil 28
17 Survey liquid chromatogram of combined fractions
from vacuum distillation 32
vi
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FIGURES (Continued)
Number Page
18 Survey liquid chromatogram of spinning band fraction 1 .... 32
19 Survey liquid chromatogram of spinning band fraction 5 .... 33
20 v. Survey liquid chromatogram of spinning band fraction 9 .... 33
21 Survey liquid chromatogram of condenser oil vacuum
stripped without heat 34
22 Survey liquid chromatogram of 100° - 105°C organic
layer from steam distillation 34
23 Survey liquid chromatogram of 100° - 105°C aqueous phase
from steam distillation 35
24 Removal of volatiles from pyrolytic oil ........ 39
25 Extraction of oil sequentially with water at 25°C,
50°C, and 95°C 40
26 Liquid chromatogram of 25°C water extract of
pyrolytic oil ...'.... 41
27 Liquid chromatogram of pyrolytic oil after successive
extraction with water at 25°C, 50°C, and 95°C ........ 41
28 Extraction of .pyrolytic oil with sodium sulfate solution ... 42
29 Combined diisopropyl and water extraction of pyrolytic oil . . 44
30 Combined anisole and water extraction of pyrolytic oil .... 45
31 Extraction of pyrolytic oil with 2% sodium hydroxide
solution 46
32 Extraction of methylene chloride solution of pyrolytic
oil with water followed by diisopropyl ether extraction
of aqueous fraction 49
33 Extraction of methylene chloride solution of pyrolytic
oil with water followed by methylisobutyl ketone
extraction of aqueous fraction 50
34 Extraction of n-butanol solution of pyrolytic oil
with water 51
35 Aqueous batch extraction, Process No. 1 54
vii
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FIGURES (Continued)
Number Page
36 Three phase extraction, Process No. 2 57
37 Sequential organic water extraction, Process No. 3 59
38 Countercurrent extractor 62
39 Separation process No. 1A—raw oil—2 stage extraction .... 78
40 Separation process IB—vacuum stripped—2 stage extraction . . 80
41 Separation process 2A—raw oil—simultaneous extraction .... 82
42 Separation process 2B—vacuum stripped—
simultaneous extraction 84
43 Pyrolysis oil pilot plant schematic—continuous process .... 87
viii
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TABLES
Number Page
1 Properties of Pine Bark-Sawdust Feed Material . . . ..... ..... . 12
2 Properties of Wood Oils from Tech-Air 50 Dry ,
Ton/day Facility 13
3 Variation of Oil Properties over Eight Months Period ...... 14
4 Typical Properties of Wood Oils and Fuel Oils .......... 15
5 Preliminary Average Molecular Weight Determinations . . . ,,. . . .; 29
6 Hydrogenations at Moderate Pressure 36
7 Hydrogenations at Intermediate Pressure 37
8 Yields of Fractions from Water Extraction of Oil . . * 42
9 Yields from Methylene Chloride Extractions of Alkaline
Solutions of Pyrolytic Oil - . . 50
10 Yields in Final Fractions from Separation Techniques
in Figures 32 and 33 51
11 Properties of Pyrolytic Oil Sample , 53
12 Composition of Yields from Batch Water Extractions,
Process No. 1 4 56
13 Composition of Yields from Batch Three Phase
Extractions, Process No. 2 t 59
14 Composition of Yields, Process No. 3 * . . 61
15 Inputs and Yields, Process 1A 64
16 Inputs and Yields, Process IB 65
17 Inputs and Yields, Process 2A 66
18 Inputs and Yields, Process 2B 67
19 Composition of Continuous Extraction Yields .... * 68
ix
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TABLES (Continued)
Number JL^&
20 Vacuum Stripping Experiments • 69
21 Organics Eluted from Aqueous Carbon Column .=.•.. 70
22 Elution of Unstripped Oil from Activated Carbon Column ..... 71
23 Distillation Data for Water-Insoluble Oil ...--.. 72
24 Analytical Results from Batch Experiment Process, 1A ...... 74
25 TLC Solvents and Detection Reagents •„» .... 75
26 Infrared Bands ....... 76
27 Liquid Chromatography Conditions • 76
28 Input Rates to Extractor 85
29 Required Extractor Volume . ...'.. 86
30 Pilot Plant - Cost Summary 86
31 Process 1A—2 Stage Continuous Extraction—Raw Oil—
Installed Equipment Cost Summary 90
32 Process IB—2 Stage Continuous Extraction—
Vacuum Stripped Oil—Installed Equipment Cost Summary 90
33 Process 2A—Continuous, Simultaneous Extraction—
Raw Oil—Installed Equipment Cost Summary 91
34 Process 2B—Continuous, Simultaneous Extraction—
Vacuum Stripped Oil—Installed Equipment Cost Summary 91
35 Depreciation—Process LA 100
36 Depreciation—Process IB 100
37 Depreciation—Process 2A 101
38 Depreciation—Process 2B
39 Price Survey of Various Chemicals
40 Return on Investment—Summary
41 Cash Flow—Process 1-A—Case I—$0.30/lb 1Q7
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TABLES (Continued)
Number Page
42. Cash Flow—Process 1-A—Case II—$0.50/lb 108
43. Cash Flow—Process 1-B—Case I—$0.30/lb 1°9
44. Cash Flow—Process 1-B—Case II—$0.50/lb 110
45 Cash Flow-Process 2-A—Case I—$0.30/lb 1H
46 Cash Flow—Process 2-A—Case II—$0.50/lb 112
47 Cash Flow—Process 2-B—Case I— $0.30/lb 113
48 Cash Flow—Process 2-B—Case II—$0.50/lb 114
49 Minimum Selling Price per Pound to Justify Investment 115
50 Average Selling Price for Pyrolytic Oil Products 124
51 Return on Investment—Percent 125
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ACKNOWLEDGMENTS
This investigation was supported by the Municipal Environmental Research
Laboratory (MERL), U. S. Environmental Protection Agency, under Grant Numbers
R 804 416 010 and R 806 403 010. We express our appreciation to Mr. Charles J.
Rogers of the Municipal Environmental Research Laboratory for his contributions,
suggestions, and encouragement during the course of this investigation.
We express our thanks to the Tech Air Corporation for supplying us with
oil samples from their 50 dry ton/day pyrolysis facility.
xii
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SECTION 1
INTRODUCTION
Large quantities of agricultural, .forestry and municipal wastes are
produced each year in the United States. The proper utilization of these
materials is of extreme importance to the country so that they can be con-
sidered a resource rather than wastes. At the same time, the disposal and
environmental problems these wastes create would be solved. One approach
for the utilization of these materials that has received a great deal of
attention in the past several years is pyrolysis. Pyrolysis of lignocellu-
losic or cellulosic material produces char, pyrolytic oil, water containing
water-soluble organic substances, and non-condensible gases. The char is
primarily carbon and can be used as a fuel or converted to activated carbon,
to producer gas for use as a clean burning gaseous fuel or to synthesis gas
for organic synthesis. The major components of the non-condensible gases
are hydrogen carbon monoxide, carbon dioxide and methane along with minor
amounts of the other hydrocarbon gases. The gas can be utilized on site as
, a clean burning low BTU gaseous fuel. The pyrolytic oils are clean burning
with heating values approximately two-thirds the heating values of fuel oils.
There is, however, a great potential for utilizing pyrolytic oils as a
source of chemical materials for industrial applications and/or as a chemi-
cal feedstock. By upgrading the oils for uses of greater value than as a
fuel the total economic benefit from waste materials would be of greater
significance to the country. Also, the utilization of oils produced from
current waste materials as a source of chemical materials would reduce the
demand on petroleum materials for chemical feedstock. In order to realize
the potential of pyrolytic oils as a source of materials for chemical appli-
cations, it is necessary to develop the processing technology to produce
refined fractions for industrial use.
Pyrolytic oils are complex mixtures of organic compounds ranging from
very volatile to high boiling materials. Many of the components are oxygen-
ated, and the oils therefore are quite different in their chemical and
physical properties from petroleum and its products. Experimental data indi-
cate that the *soil may contain as many as 200 or more compounds. The charac-
terization of the pyrolytic oils as produced and fractions obtained from
them by determination of physical and chemical properties provides data
needed for the development of the technology to process the oils into more
useful chemical materials.
The overall approach to developing technology for processing the oils
to yield more useful fractions has been mainly with distillation techniques
and separation (extraction) techniques. Distillation experiments include
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atmospheric and vacuum distillation, fractional distillation, steam distil-
lation and vacuum stripping of water and volatile components. Separation
techniques include extraction with water at different temperatures and an
aqueous salt solution, simultaneous extraction with water and an organic
solvent, extraction with alkaline solutions, and extraction of organic
solvent solutions of pyrolytic oils with water. The extraction techniques
show promise of having the greatest potential for processing the oil into
fractions containing fairly specific chemical classes of compounds. These
fractions should find ready utilization in industrial applications. Distil-
lation offers more promise as a method for processing a specific fraction
of oil into more highly refined and purified products. There are several
potential approaches utilizing extraction techniques which could produce
three or four oil fractions that would have potential for industrial appli-
cations. Experimental work was conducted at the bench level on both a bath
basis and a continuous basis. Based on the results from the continuous
extraction experiments, a pilot plant has been designed for investigating
the continuous processing of pyrolytic oils. Also, the preliminary economics
of processing the pyrolytic oils on a commercial scale have been evaluated.
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SECTION 2
SUMMARY
Oils produced by the Georgia Tech pyrolysls process from the Tech-Air
50 dry ton/day pyrolysis facility have been thoroughly characterized. The
overall chemical and physical properties have been determined by standard
analytical techniques. The oils are dark brown to black and have a burnt,
pungent odor. The viscosity of the oils depends upon a number of factors,
such as the pyrolysis mode, the operating conditions and the amount of water
emulsified in the oil. Oils which contain 10 to^ 15% water are relatively
free flowing. The oils have heating values which are approximately two-thirds
the heating values of petroleum fuel oils and burn cleanly. The oils are
acidic and exhibit some corrosive characteristics.
The oils are complex chemical materials with a wide spectrum of oxy-
genated compounds which exhibit a variety of functional groups and wide
boiling range. The chemical composition of the oils is of importance in
devising processing methods for producing useful chemical fractions from the
raw oils. The analytical techniques of choice for determining the chemical
composition of the oils and fractions produced from them are liquid chro-
matography, thin layer chromatography, gas chromatography and gas chromato-
graphy/mass spectroscopy. The major classes of organic chemical species
found in the pyrolytic oils investigated in this program from forestry
materials were phenolics, polyhydroxy neutral compounds, neutral compounds of
high aromaticity, and volatile acidic compounds.
The development of processing methods to produce fractions of the oils
for potential chemical applications was focused on producing fractions which
would contain predominantly a specific class of compounds. Distillation is
a highly developed chemical operation and offers a possible method for
processing and refining pyrolytic oils. Therefore, various distillation
techniques were tested. Due to the heat sensitivity of the oils, the pyro-
lytic oil in the flask, after distilling about 50% to 65% of the charge,
would begin to decompose. In addition, fractional distillation at low
pressure did not produce any narrow cut fractions over the whole boiling
range with a predominantly chemical species. Although distillation should
not be considered as the initial processing step for pyrolytic oils, it
should be considered as means of refining fractions of the oil produced by
other processes. Some preliminary catalytic hydrogenations were carried out
at about four and 20 atmospheres pressure. Based on the results of these
experiments, hydrogenation should not be considered as the first processing
step, but should be considered as a potential refining method for some of
oil fractions produced by other processing methods.
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Separation processes based on extraction techniques employing the
solubility of the oil in water and various organic solvents .offer a poten-
tial approach for separation of the oil into three or four major fractions,
each of which would contain a predominant chemical species. Five major
approaches involving extraction techniques were tested at the bench level on
a batch basis. These approaches were extraction: (1) with water at differ-
ent temperatures; (2) with sodium sulfate solution (salting-out effect);
(3) with water and a water-insoluble organic solvent "(three phase system);
(4) of sodium hydroxide solutions at different pH ranges with methylene
chloride; and (5) of organic solvent solutions of oil with water. The
results of these extraction techniques and experiments showed promise and
the approaches selected for additional work at the batch level were aqueous
extraction, simultaneous extraction with waiter and an organic solvent and
aqueous extraction of an organic solution of the oil!. Both vacuum stripped
and unstripped oil samples were examined by all three processes and the
effects of both polar and nonpolar solvents were studied.
Based on the results of these experiments, aqueous extraction (Process
No. 1) and simultaneous extraction with water' and an organic"solvent (Process
No. 2) were selected for continuous extraction experiments at the bench
level. Continuous extraction experiments were conducted with both vacuum
stripped and unstripped oil samples. The data from the continuous experi-
ments indicated the complexity of processing pyrolytic oils. The oils have
a large number of compounds which exhibit a wide boiling point range and a
high degree of chemical functionality and chemical nature, such as solubility,
polarity, etc. The results from the continuous experiments show that both
aqueous extraction, Process 1, and simultaneous extraction, Process 2 , have
promise as the initial steps in processing pyrolytic oils. The insoluble
oil phases from Process 1 and the MIBK phases from Process 2 did not contain
any polyhydroxy neutral compounds, based on the analysis. The aqueous phases
from both Processes 1 and 2 contained phenolic, polyhydroxy neutral com-
pounds, and neutrals of high aromaticity. MIBK extraction of these aqueous
phases removed the major portion of the neutrals of high aromaticity. Pre-
liminary extraction experiments with alkali solution of the MIBK fractions
showed that the phenolic fraction could be removed, which would provide two
fractions, one predominantly phenolics and the other predominantly neutrals
of high aromaticity. In order to obtain fractions of the oils which contain
predominantly a group of compounds that are chemically similar, it would be
necessary to further process the phases obtained by extraction techniques.
Additional processing could include extraction steps and distillation.
In order to produce fractions of oil for chemical applications from
raw pyrolytic oil from biomass, there are two major areas that need further
investigation. Additional experimental work must be conducted at the small
scale pilot plant level to yield suitable fractions of the oils for inves-
tigations for industrial applications and to produce data for the design of
a commercial plant. In addition, the studies at the pilot plant level should
include additional processing, such as distillation, of the fractions obtained
by the extraction techniques. The application studies for the oil studies are
necessary as each fraction would consist of a mixture of compounds. A versa-
tile pilot plant was designed for testing at the rate of four gallons per
minute, the extraction processes developed in this program. Additional
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processing of the fractions, such as distillation, could also be investigated
with the pilot plant. The processing of pyrolytic oils with the pilot plant
could be optimized to produce fractions most suitable for industrial uses as
indicated by application studies and to provide the data for design of a com-
mercial plant.
Preliminary economic assessments of the processing of pyrolytic oils
were made, based on two approaches. These preliminary assessments are prom-
ising, lii one approach, the average selling price per pound for the processed
oil products was determined that would be necessary to provide a 15, 30 and 50
percent net return on investment. For a 50 percent return, the price range of
8.4 to 10.6 cents per pound is in the same range as 9 cents per pound for coal
tar creosote and well below 54 cents per pound for coal tar cresylic acid,
which were quoted market prices in December, 1979. In the other approach,
two schedules of selling prices were assumed for each product in Processes 1
and 2, based on quoted market values of chemical materials which were consi-
dered to be similar. The returns on investment were very promising for both
price schedules. The significance of this economic assessment is that at a
relative low selling price, processing of pyrolytic oils should be economi-
cally viable and that if suitable industrial applications for the processed
oil fractions can be found, processing pyrolytic oils should be very profi-
table.
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SECTION 3
RECOMMENDATIONS
The results of this study have indicated that the processing of pyrolytic
oils from wood into products suitable for commercial applications is technically
feasible and the preliminary economic assessment is very promising. However,
additional research and development work is needed so that this industrial
potential for pyrolytic oils can be realized. The two major areas in which
additional work is required are processing studies with pyrolytic oils at the
pilot plant level and studies on utilization of the products in industrial'
applications.
It is recommended that investigations with pyrolytic oils be conducted
at the pilot plant level with both aqueous extraction (Process 1) and simultan-
eous extraction with water and an organic solvent (Process 2). With both pro-
cesses, additional processing of the initial phases should be investigated, and
both raw oil and vacuum stripped oil should be tested. The objectives of this
program would be to develop optimum operating conditions for producing suitable
oil fractions for industrial applications, to obtain engineering data for scale
up for a commercial plant, to produce sufficient quantities of oil fractions to
use in a study for industrial utilization, and to obtain adequate data to -make
an economic analysis of the process and of the potential market for the products.
A significant part of these recommendations is the investigation for potential
chemical applications for the oil fractions, such as utilization in the produc-
tion of resins. The objective of this phase of the program would be to estab-
lish specific applications for the oil fractions and to determine the potential
markets. The results of this recommended program should provide the necessary
information and data for the utilization of pyrolytic oils in chemical appli-
cations on an industrial scale.
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SECTION 4
BACKGROUND INFORMATION
PYROLYSIS AND DESTRUCTIVE DISTILLATION OF WOOD
Pyrolysis is an old process and has been used industrially in the past
on'a batch basis to produce charcoal, pyroligneous liquor (mostly water with
dissolved organic compounds), insoluble tars, and non-condensible gases.
It was utilized during and after World War I in this country and was known
as wood distillation. With the utilization of petroleum as a chemical feed-
stock, the pyrolysis process became uneconomical and is no longer practiced
in this country. Various aspects of destructive wood distillation and the
products have been discussed in representative literature references [1, 2,
3, 4 and 5].
f~
The destructive distillation of wood was generally carried out as a
batch process in a retort with external heat and produced the products men-
tioned above. The significant and important difference between the Engi-
neering Experiment Station pyrolysis process and the old wood distillation
process is that the Engineering Experiment Station pyrolysis process is a
self-sustained continuous process. This is of significance because the
pyrolytic oil produced in this manner from a given feed material under
specific operating conditions is a reproducible product with definite
physical and chemical properties. Therefore, it has potential as a feed-
stock for processing into other products on a commercial scale. Its poten-
tial for uses other than as a fuel warrants extensive investigation.
GEORGIA TECH PYROLYSIS PROCESS
The Georgia Tech pyrolysis process is a continuous, self-sustained
pyrolysis system which was developed over the past several years by staff
members of the Engineering Experiment Station. Particular attention is
devoted to this process since all the pyrolytic oil used in this investiga-
tion was produced in either one of the pilot plants on the Georgia Tech
campus or at the field development facility owned by the Tech-Air Corporation.
A wide variety of agricultural, forestry and municipal wastes have been pro-
cessed under a variety of operational conditions with the Engineering
Experiment Station pilot plant pyrolysis systems.
jg
Licensed to the Tech-Air Corporation, a wholly owned subsidiary of the
American Can Company.
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Background Experience and Pilot Plants - Georgia Tech Pyrolysis Process
Workers at the Engineering Experiment Station, Georgia Tech, have
found that pyrolysis is readily adaptable for the conversion of cellulosic
and lignocellulosic wastes into useful fuels and other products. Involve-
ment at Georgia Tech in the area of conversion of solid wastes by pyrolysis
began with work in 1968 to develop a means to dispose of peanut hulls with-
out producing the pollution problems of incineration.
The steady-flow, low temperature pyrolysis process developed at the
EES involves processing of the wastes in a porous, vertical bed. Among the
advantages of the process are its simplicity and its low temperature opera-
tion. These features, together, lead to a highly economical design. In
addition, the system is self-sustaining and requires a minimum of processing
of the wood wastes prior to pyrolysis [6,7].
The first pilot plant system, approximately five feet tall, was
designed to reduce peanut hulls to a char and a combustible gas. The system
built in 1968 was operated on a batch basis at first and then on a continu-
ous basis with a manual input feed. Hundreds of pounds of peanut hulls were
converted to char and off-gases during several months of testing with this
equipment. Enough data were obtained to demonstrate the feasibility of
developing an automated prototype converter with the vertical, porous bed
design.
The large prototype, constructed in 1971, was built to operate contin-
uously at an input feed rate of 4,000 pounds per hour. The unit was approxi-
mately 11 feet in height, and the reaction chamber was mounted on top of a
water-cooled collection chamber. The feed-out was accomplished by a hori-
zontal screw at the base of the chamber. The off-gases were treated as <„*.
potentially explosive in these tests, and consequently, a system was con-
structed to burn the gases in an unconfined, diffusion controlled flame.
Experience with these gases showed that they could be burned safely and
easily by premixing and igniting in a conventional fashion. This system was
operated over a period of many months, while processing thousands of pounds
of feed. The reaction chamber of this converter was designed to have a mini-
mum weight and only enough operating life to demonstrate the automatic
operation of the process. This was done to reduce the overall cost of this
experimental prototype. Consequently, the test program started with low
temperature operation and on succeeding tests the temperature was raised.
The internal structure of the reaction chamber eventually failed after
approximately six months of testing as a result of the elevated temperature.
Based on the data and results from the first pilot plant unit and the
experimental prototype, a third pilot plant was designed and built. This
system was used to process a wide variety of feed materials to determine
operating characteristics and investigate operating parameters. This system
was completely rebuilt in the fall of 1975. Presently, the system includes
a waste receiving bin, a belt conveyor to the converter, the converter and
char handling system, an off-gas cyclone, a condenser by-pass, demister,
draft fan, and vortex after-burner. The present system will process 500 to
800 Ibs. waste/hour depending on the density of the feed material. Types of
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waste processed through the converter include peanut hulls, wood chips, pine
bark and sawdust, automobile wastes, municipal wastes, macadamia nut shells,
and cotton gin wastes. The pyrolytic oil used in the third phase of the
experimental program was produced in this unit in 1978.
The fourth Engineering Experiment Station pyrolysis pilot plant, which
is larger and more versatile, was designed, assembled and put into operation
by the staff of the Engineering Experiment Station in September, 1974. This
unit has a design capacity of 1,500 pounds of dry material/hour and has been
used extensively to test municipal wastes, peanut hulls, and wood wastes.
Commercialization of EES Pyrolysis Process
The pyrolysis process developed by workers of the Engineering Experi-
ment Station, Georgia Tech, was licensed to the Tech-Air Corporation in 1971
for commercialization. Tech-Air field tested pyrolysis converters at a
peanut shelling plant and a lumber yard. The most extensive field testing
and development,program was conducted at a lumber yard in Cordele, Georgia,
over a five year period. The Tech-Air field demonstration facility processed
approximately 40 dry tons/day of a mixture of pine bark and sawdust and pro-
duced char, oil and noncondensed gases. The char was used for making char-
coal briquettes, the oil was sold as a fuel, and the gases were being used
on-site as a fuel for drying input feedstock. The char and oil can be
stored and transported, and the noncondensed gases must be burned on-site.
In the Tech-Air demonstration facility part of the combustion gases are used
in a drier of Tech-Air design to reduce the moisture content of the feed
material to less than 10%. The input feed material varies in moisture con-
tent from 30% to 55% on a wet basis, depending on weather conditions, season
of year, and amount of sawdust in the feed. A number of improvements were
made in the system, and the system was operated for a period of several
months on a 24 hour basis with a reliability of operation at design through-
put of "better than 90%. An analysis of the combustion stack gases was made
and comparison of these data with the EPA exhaust standards revealed that
the system easily met all federal standards. The Georgia Tech pyrolysis
system can be operated in a highly reliable manner with a wide range of
feed materials and offers a high degree of flexibility for the conversion
of agricultural and forestry residues and municipal wastes to char, oil and
gas. The pyrolytic oil for the first and second phases of the experimental
program was produced in this facility.
PYROLYTIC OIL FROM WASTE MATERIALS
Pyrolytic oil from different waste materials represents a potential
source of feedstock for the chemical industry and/or as a source of chemicals.
It has been reported that about six percent of United States consumption for
oil goes for feedstock for the chemical industry [8]. On an annual basis
this would amount to approximately 50,000,000 tons of petroleum. The yield
of pyrolytic oil from lignocellulosic material processed by the Engineering
Experiment Station pyrolysis process varies from 15 to 25 percent depending
upon feed material and operation conditions. Consequently, it would require
-------
200 to 330 million tons of dry lignocellulosic material to supply a tonnage
of pyrolytic oil in the same tonnage range of petroleum used by the chemical
industry. It should be pointed out that this does not imply that pyrolytic
oil would be processed in the same manner as petroleum feedstock or that one
ton of pyrolytic oil is equivalent on a feedstock basis to one ton of
petroleum.
Accurate estimates of wastes from different sources are difficult to
obtain. Based on our inquiries, particularly with the U. S. Forest Service,
the amount of forestry wastes in the U. S. is estimated at 100 million dry
tons annually (Heywood T. Taylor, U.S.F.S., Private Communication). This
quantity of material has the potential of supplying 33 to 50 percent of the
tonnage of petroleum now used by the chemical industry. The significance of
these data is that from the standpoint of quantity the potential exists for
pyrolytic oil from forestry wastes alone to make a significant contribution
as a source of chemical feedstock. Anderson in 1972 estimated in his study
net oil potential of 1.1 billion barrels of oil per year from the total
organic wastes generated annually in the U. S. [9]. Tillman has recently
reported that there is a potential source of approximately one billion dry
tons of cull or rough trees and salvable dead trees in the U. S. [10]. The
important fact that these data provide is that there are large quantities of
waste material which have the potential for being converted to resources,
and therefore, making a real impact on the material and energy needs of the
U. S.
10
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SECTION 5
EXPERIMENTAL—PHASE I
ANALYSIS AND CHARACTERIZATION OF PYROLYTIC OILS
The oils obtained from the pyrolysis of lignocellulosic materials are
complex mixtures of organic compounds and usually contain some water. Con-
sequently, the characterization of the physical and chemical properties of
pyrolytic oils requires that one use a variety of analytical and testing
techniques. Properties that are of interest in characterizing pyrolytic oils
include but are not necessarily limited to density, water content, heating
value, acidity, flash point, pour point, corrosiveness, filterable solids,
ash, solubility in various solvents, distillation range, viscosity and ele-
mental content, particularly carbon, hydrogen, nitrogen, sulfur and oxygen.
The identification of the chemical species and compounds and the rela-
tive quantities are data that are needed for developing methods for utiliza-
tion of the oils for applications other than as a fuel oil. Among the most
useful techniques for obtaining this information and data are gas, thin-layer
and liquid chromatography, gas chromatography/mass spectroscopy, and infrared
and ultraviolet spectroscopy.
Sources of Oil
Samples of pyrolytic oils for Phase I were obtained from two major
sources: (1) the 50 dry tons/day field demonstration pyrolysis facility of
the Tech-Air Corporation at Cordele, Georgia, and (2) the 500 to 800 Ibs/hr
pyrolysis pilot plant (Blue IV) of the Engineering Experiment Station,
Georgia Tech, which is operated on campus. Some samples of oil were produced
in a six inch tube furnace fitted with a condensation train and gas collec-
tion system. A complete description of this apparatus and the pyrolysis pro-
cedure has been reported
The physical and chemical characteristics of pyrolytic oils depend
upon the feed material, the pyrolysis process and the conditions under which
pyrolysis occurs. In the old wood distillation industry, the retort batch
process produced organic materials which varied from the low boiling com-
pounds such as methyl alcohol to the insoluble tars. Continuous pyrolysis
processes of today, such as the Georgia Tech process [6, 7], can be oper-
ated at steady state conditions with a given feed material to produce oils
of fairly constant compositions and properties. These oils have greater
potential than those from the old wood distillation industry as a source of
chemical materials for industrial applications and are much more suitable
feedstock for continuous processing to produce fractions of oil suitable for
11
-------
specific applications. For these reasons, the oils used in this investiga-
tion were mainly those produced in the continuous pyrolysis facility of the
Tech-Air Corporation or in the pyrolysis pilot plant of the EES, Georgia
Tech.
Samples of oil were obtained from the Tech-Air facility in July, 1976,
and May, 1977. In each case, oil samples were obtained from the air-cooled
condenser and the draft fan. The feed material for this facility was pine
bark-sawdust, and a representative sample had the properties listed in TABLE
1.
TABLE 1. PROPERTIES OP PINE BARK-SAWDUST FEED MATERIAL
Property Results Method
Pinebark 70 Microseparation by
Pine sawdust 30 visual means
Bulk density 213 kg/m „
(13.3 Ibs/ft )
Moisture 10.3% ASTM D-1762-64' ••. '
Ash (weight %) 1.3% ASTM D-1762-64
Acid Insoluble <0.1% - £.. , (.
Ash (weight %)
Heating Value 21.2 MJ/kg ASTM D-240-74
(dry basis) (9109 Btu/lb)
Oil samples, produced in the Georgia Tech pilot plant on July 22 and
27, 1977, from pine chips and on September 16, 1977, from hardwood chips,
were also used in these studies.
During the course of this investigation, samples of oil have been sup-
plied to Dr. M. B. Polk of Atlanta University for use on E.P.A. Grant No.
R 804 440 010. The oil samples provided were those obtained from Tech-Air in
July, 1976, and May, 1977, and those produced in the Georgia Tech pilot
plant in July, 1977, from pine chips and in September, 1977, from hardwood
chips. In addition, oil samples produced in the six inch tube furnace
pyrolysis facility (batch process) from a pine bark-sawdust mixture and
hardwood chips were supplied.
Analytical and Test Data—
The condenser and draft fan oils obtained from the Tech Air facility in
July, 1976, were characterized extensively, and the results are illustrative
of the physical and chemical properties of pyrolytic oils and of the many
analytical techniques and methods that can be used [12] . The data for the
condenser and draft fan oils from the Tech-Air ton/day facility are given in
TABLE 2.
12
-------
TABLE 2. PROPERTIES OF WOOD OILS FROM TECH-AIR 50 DRY TON/DAY FACILITY
Property
Density
Water content
(weight %)
Heating Value
(wet basis)
PH
Acid Number
Flash Point
Filterable Solids
(weight %)
Copper Strip
Corrosion
Sulfur (weight %)
Pour Point
Ash (weight %)
Distillation
First Drop
10% Point
48% Endpoint
53% Endpoint
Solubility
(weight %)
Acetone
Methylene
Chloride
Toluene
Hexane
Elemental Analysis
(weight %)
Carbon
Hydrogen
Nitrogen
Condenser Oil
1,141 kg/m3
(9.525 Ibs/gal)
14.0%
21.2 MJ/kg
(9,100 Btu/lb)
2.9
75 mg KOH/g
111°F
(233°F)
0.3%
1
0.01%
26.7°C
(80°F)
0.08%
98°C
103°C
NA
282°C
99.6%
93.5%
Slightly
Slightly
51.2
7.6
0.8
Draft Fan Oil
1,107 kg/m3
(9.242 Ibs/gal)
10.4%
24.6 MJ/kg
(10,590 Btu/lb)
3.3
31 mg KOH/g
121°C
(240°F)
0.4%
1
0.01%
26.7°C
(80°F)
0.03%
101°C
105°C
265°C
NA
99.6%
97.8%
Slightly
Slightly
65.6
7.8
0-9
Method
-
ASTM D 95-70
ASTM D 240-64
5% Oil dispersed
in water
ASTM D-664-58
ASTM D-93-73
Acetone Insoluble
Classification-
ASTM D-130-7
ASTM D-129-64
ASTM D-97-66
-
ASTM D-86
Group 3
-
-
13
-------
Samples of the condenser and draft fan oils were stored at ambient
temperature and 0°C for approximately eight months and then certain proper-
ties were determined. These data, presented in TABLE 3, show that the oils
can be stored for periods of five to six months without any deleterious
effects if the oils are to be used as fuels only. If the oils are to be
used as a source of chemical materials, then it would be necessary to con-
sider the effect of storage on the processing characteristics of the oils.
TABLE 3. VARIATION OF OIL PROPERTIES OVER EIGHT MONTHS PERIOD
Property
Stored Eight Months
Initial Value
0°C
Ambient Temperature
Condenser Oil
Water Content
(weight %)
Heating Value
(wet basis)
Acid Number
Viscosity*
pH
Water Content
(weight %)
Heating Value
(wet basis)
14.0%
21.2 MJ/kg
(9,100 Btu/lb)
75 mg KOH/g
0.275 Pa
2.6
20.5%
22.8 MJ/kg
(9,800 Btu/lb)
87 mg KOH/g
0.350 Pa
3.4
Draft Fan Oil
10.4%
24.6 MJ/kg
(10,590 Btu/lb)
15.5%
24.8 MJ/kg
(10,660 Btu/lb)
24.1%
21.4 MJ/kg
(9,190 Btu/lb)
89 mg KOH/g
0.175 Pa
2.9
12.7%
24.9 MJ/kg
(10,690 Btu/lb)
Acid Number
Viscosity*
PH
31 mg KOH/g
0.233 Pa
3.3
71 mg KOH/g
0.079 Pa
3.1
60 mg KOH/g
0.475 Pa
3.0
* Determined with Brookfield Viscosimeter, Model LV with Thermosel system at
25°C at 60 r/min.
Some typical properties of the condenser and draft fan oils and fuel
oils are compared in TABLE 4.
Viscosity—The viscosity of liquids and its change with temperature
is a significant property in the material handling and processing of liquids.
A Brookfield viscosimeter, Model LV, with Thermosel system was used to deter-
mine viscosity values. The viscosity versus temperature was determined for
both the condenser and draft fan oils initially and on samples which had
14
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TABLE 4. TYPICAL PROPERTIES OF WOOD OILS AND FUEL OILS
Wood Oils*
Fuel Oils
Property
Water Content, %
Heating Value, MJ/kg
(Btu/lb)
(Btu/gal)
3
Density, kg/m
(lb/gal)
Pour Point, °C
Flash Point, °C
Viscosity, Pa's*
Elemental Analysis
Carbon %
Hydrogen %
Nitrogen %
Sulfur %
Condenser
14
21.2
(9,100)
(86,700)
1,141
9.525
26.7
111
0.225
51.2
7.6
0.8
<0.01
Draft Fan
10.4
24.6
(10,590)
(97,850)
1,107
9.242
26.7
121
0.233
65.6
7.8
0.9
<0.01
#2
Trace
45.7
(19,630)
(139,400)
851
7.10
-18 max
38 min
0.020
86.1
13.2
0.6-0.8
#6
2
43.2
(18,590)
(148,900)
960
8.01
18-29
65
2.262
87.0
11.7
0.9-2.3
* Values obtained on oils with moisture content as reported.
' Values for fuel oils are considered typical. Sulfur will vary depending
on origin of oil. Ref., North American Combustion Handbook, 1st ed.,
North American Mfg. Co., Cleveland, Ohio, 1952.
| Determined with Brookfield Viscosimeter,' Model LV with Thermosel system at
25°Ciat 60 r/min.
been stored at 0°C and ambient temperature for approximately eight months.
These viscosity curves are given in Figures 1 and 2. The viscosity versus
temperature curves of samples of both oils which had been vacuum stripped for
removal of water and volatiles are given in Figures 3 and 4. In order to
determine the effect of prolonged heat upon the viscosity of condenser oil,
samples of sealed oil were heated at 110°C for different time periods, and
the viscosity was then determined for each sample. These data are presented
in Figure 5. For comparison, the viscosities of the condenser oil and #2
and #6 fuel oils are presented in Figure 6.
Liquid chromatography—The wood oils are heat sensitive, reactive and
contain a relatively large number of organic compounds. An analytical tech-
nique was needed which could be used in analyzing the fractions of oil
obtained by the different processing methods that would not change the chem-
-ical character of the fractions. Liquid chromatography (LC) appears to be
the method of choice because LC is carried out at ambient temperature, is
15
-------
Cd
w
M
O
PM
1-1
H
O
A - Initial viscosity curve
B - Sample stored at 0° C for eight months
C - Sample stored at ambient temperature
for eight months
_L
_L
20 30 **Q 50 60 70 80 90 100
TEMPERATURE, °C
Figure 1. Viscosity of condenser oil.
600
500
w
en 400
H
O
1 30°
CJ
200
100
- Initial viscosity curve
- Sample stored at 0° C for eight months
- Sample stored at ambient temperature
for eight months
_L
_L
20 30 'tO 50 60 70 80 90 100
TEMPERATURE, °C
Figure 2. Viscosity of draft fan oil.
16
-------
Pd
CO
M
O
A - Initial viscosity curve
B - Vacuum stripped viscosity curve
l
20 30 kO 50 60 70 80 90 100
TEMPERATURE, °C
Figure 3. Condenser oil.
17
-------
rn
UJ
u
1900
1800
1700
1600
1500
1400
1300
1200
1100
1000
900
800
700
600
500
400
300
200
100
A - Initial viscosity curve
B - Vacuum stripped viscosity curve
20 30 40 50 60 70 80 90 100
TEMPERATURE, °C
Figure 4. Draft fan oil.
18
-------
M
O
W
CJ
1800
1700
1600
1500
11+00
1300
1200
1100
1000
900
800
700
600
500
400
300
200
100
20 30 40 50 60 70 80 90 100
TEMPERATURE, °C
Figure 5. Effect of heating condenser oil at 110°C
for different time periods on viscosity.
19
-------
1900
1800
1700
1600
1500
11*00
1300
1200
1100
1000
900
800
700
600
500
, 400
300
200
100
0
- Condenser oil
- No. 2 Fuel oil
- No. 6 Fuel oil
20
30 kO 50 60 70 80 90 100
Figure 6. Viscosity curves for condenser oil (initial)
and No. 2 and No. 6 fuel oils.
20
-------
capable of high resolution of complex mixtures, and component detection is
nondestructive. In addition, the wood oils are soluble in organic-aqueous
solvent systems which are very useful in LC. The main initial objective of
utilizing LC in the work with the wood oils is to provide a method to obtain
"fingerprints" of the raw oil and fractions produced from it for comparison
and correlation.
Testing of LC Variables—
The variables that were studied to find satisfactory LC conditions
were LC columns, uv wave length, solvent gradient and solvent flow rate. The
condenser oil (July, 1976) was used for testing all of these variables.
LC columns—In order to select the most suitable LC column, several
columns were tested with the raw condenser wood oil (July, 1976) using one
ml/min flow rate and uv detector at 254 nm. The chromatographic columns and
conditions tested and the results are given below in the order in which the
testing was carried out.
A. Vydac adsorption silica gel 30y column. Solvent, 0-100%
2-propanol in isooctane, 20 min gradient 20 concave.*
Results: No resolution obtained; only one large peak.
B. Partisil adsorption silica gel 5p column. Solvent, 5-30%
2-propanol in isooctane, 20 min gradient, linear. Results:
Resolution of only eight peaks.
C. Partisil PAC 5\i column. Solvent, 0-100% 2-propanol in iso-
octane, 30 min gradient 35 concave. Results: Resolution
of 12 to 20 peaks. See Figure 7.
D. Partisil ODS 5y column. Solvent, 10-100% acetonitrile in
water, 30 min gradient 35 concave. Results: Resolution
of 30-40 peaks. See Figure 8.
E. Partisil ODS 5y column. Solvent, 10-100% acetonitrile in
water, 10 to 40% with 20 minute hold, then 40% to 100% 35
concave gradient. Results: Resolution of 46-50 peaks.
Total run time 60 minutes. See Figure 9.
F. Partisil ODS 5y column. Solvent, 10-100% acetonitrile in
water, 30 min linear gradient. Results: Better overall
presentation of chromatogram and better resolution of later
peaks without excessive runtime.
From the above results, the resolution obtained with the conditions given in
D above are very suitable for our survey chromatograms and the conditions
in E and F for obtaining of greater resolution.
Wavelength—The wavelengths 200, 220, 254, 280, 300, 320, 360 nm
were selected and LC runs were made using constant conditions (E above)
other than wavelength. The results were: (a) It was noted that many compo-
nent responses appeared or disappeared with the change in wavelength;
(b) no one wavelength was entirely satisfactory because at the shorter wave-
lengths of 200-220 nm peak resolution; (c) the longer wavelength of 300-360
nm produced sharply resolved peaks, but only a small total number of peaks
*Term used as a dial setting for logarithmic slope control on Micrometritics
LC models only.
21
-------
Figure 7. Liquid chromatogram of wood oil. Partisil PAC column with 0-100%
solvent gradient of 2-propanol in iso-octane.
Figure 8^ Liquid chromatogram of wood oil. Partisil ODS column with 10-100%
solvent gradient of acetonitrile in water.
22
-------
NJ
U)
cu TJ
in ft>
0
N3«
CTJ C
rT1 —:
/=k -""
--• r? &>
li-g
:q^
Figure 9. Liquid chromatogram of wood oil. Partisil ODS column with 10-100% solvent
gradient of acetonitrile in water with 20 minute hold at 40% acetonitrile.
cr
53
-------
actually appeared; (d) and the most satisfactory results for our purposes
were obtained at 280 nm with 254 nm being the alternative choice. See
Figures 10 through 14 for representative liquid chromatograms of this study
with condenser wood oil using conditions in E above.
Solvent gradient—The water-acetonitrile solvent system was found to
be satisfactory for these wood oils. Water-methanol was tested but was
unsatisfactory.
A. 10-100% acetonitrile solvent gradient with 35 concave
instrument setting, 30 min long run with no solvent holds
produced a short, fairly well resolved chromatogram with
crowding of peaks during the last 25% of the run. See
Figure 8.
B. A 10-40% acetonitrile solvent gradient with 35 concave
instrument setting, and solvent hold for 20 min, then to
100% for 10 min produced a very well resolved chromatogram
in 60 min. This run produces typically 50 discernible peaks
from the raw condenser oil test sample. See Figure 9.
C. A gradient with 5 min solvent holds at 20%, 30%, 40%, then
10 min at 100% did not produce a better resolved chromatogram
than condition B. Condition B was selected as a standard
gradient with condition A being used for survey scans.
Flow rate—Liquid chromatograms were made using flow rates of 1 ml/min,
2 ml/min and 0.5 ml/min. A flow rate of 1 ml/min was selected because it
produced the best resolution consistent with a practical time limitation of
1 hour per LC run.
Liquid Chromatograms of Wood Oils—
Two sets of liquid chromatographic conditions were selected for obtain-
ing liquid chromatograms of the oil samples. Survey liquid chromatograms
are obtained with the conditions given in D and greater resolution liquid
chromatograms are obtained with the conditions given in E in the above dis-
cussion on liquid chromatography. Survey liquid chromatograms are presented
in Figures 15 and 16 for the condenser and draft fan wood oils obtained
July, 1976, from the Tech-Air Corporation. An examination of these chroma-
tograms shows that all of the samples have a large number of components and
that each chromatogram has distinctive features.
Molecular Weight Determinations of Oils by LC—
The results from the processing of wood oils from pyrolysis of wood,
particularly when subjected to heat, indicate that reactions occur which
produce higher molecular weight components. It is also desirable to have
information on the molecular weight distribution of the raw wood oils. In
an attempt to obtain some information which would be indicative of the
molecular weight range of the oils and fractions of oil, the newly available
size exclusion liquid chromatographic columns of silica gel with narrow
pore size distribution were utilized. The column selected was a 25 cm col-
umn of DuPont SE-60 controlled size deactivated silica which has a molecular
weight range of linear operation of approximately 100 to 800 Mw. Polystyrene
standards of 800, 2200 and 9000 were obtained from Pressure Chemical Company,
24
-------
Ln
Figure 10. Liquid chromatogram of wood oil at 210 nm.
Figure 11. Liquid chromatogram of wood oil at 254 nm.
-------
Figure 12. Liquid chromatogram of wood oil at 280 nm.
Figure 13. Liquid chromatogram of wood oil at 300 nm.
-------
l-o
Figure 14. Liquid chromatogram of wood oil at 360 run.
-------
20
Minutes
25
30
35
Figure 15. Survey liquid chromatogram of raw condenser oil.
10
15
20
Minutes
25
30
35
Figure 16. Survey liquid chromatogram of draft fan oil.
28
-------
Pittsburgh, Pennsylvania. Benzene, molecular weight 78, was also used. In
these LC runs, the solvent was tetrahydrofuran and the UV detector was set
at 280 nm. The average molecular weights of raw wood oils and some oil frac-
tions were obtained. In addition, the still bottoms from a commercial
distillation of a wood oil was tested. The preliminary results from this
initial work are given in TABLE 5.
TABLE 5. PRELIMINARY AVERAGE MOLECULAR WEIGHT DETERMINATIONS
Sample Description Mw Comment
Raw Condenser Oil 160
Raw Draft Fan Oil 150
Still Bottoms from Atmospheric 150
Distilled Oil
Vacuum Spinning Distillation
Fractions 1-4 (combined) 100
Fraction 8 80 and 120 Two Main Components
Fraction 12s 150
Still Bottoms Steam Distilled Oil 150
^3to»
Still Bottoms "from Commercially
Distilled Oil* 112 - 9000 Broad Mw Distribution
.^ -rip
* Obtained from Tech-Air Corporation
Gas Chromatography
Gas chromatography (GC) offers an excellent technique for analyzing
complex mixtures of organic compounds. The apparent disadvantage in
analyzing wood oils (produced by pyrolysis) by GC is the heat sensitivity of
some components in wood oils and the possible effect of the heat on these
components during GC analysis. Recognizing this possible constraint, GC
should be useful for analysis for fractions containing more volatile compo-
nents, for water soluble components and for fractions obtained in experiments
designed to separate pyrolytic oils into fractions containing a major chemi-
cal class of compounds.
In addition, it was considered appropriate to do some preliminary
analysis of the raw wood oils because of the powerful analytical capability
of GC. The instruments used were a Perkin Elmer Model 900 with a flame
ionization detector with dual column and temperature programmed capability,
and a Perkin Elmer Model 990 with thermal conductivity detector, dual column,
and isothermal oven.
29
-------
The objectives of this gas chromatographic work are to be able to
resolve the low molecular weight components in the aqueous phases of various
distilled fractions, to resolve the more volatile components of the oils and
fractions of oil, and to analyze the higher molecular weight components of
the relatively water-free wood oils and fractions obtained from the oils.
To date, two columns were selected from several GC trial runs with the raw
condenser oil and a distilled aqueous fraction. The list of columns and con-
ditions that have been tried are given below.
Initial Conditions: P.E. 900 FID detector. Carrier gas, N? at 20
ml/min temperature program as shown.
P.E. 990 T.C.. detector. Helium carrier gas at 20 ml/min;
isothermal oven.
Samples tested were raw condenser oil and aqueous distilla-
tion fraction.
Column 1. Porapak Q, 9' x 1/8", with 1' x 1/8" Porapak Q precolumn
to retain and prevent the heavy organics from entering the
main column. Oven 120°C, injector 200°C, thermal conduc-
tivity 225 ma, Helium carrier at 20 ml/min. Results: The
determination of water, lower alcohols, formaldehyde and
acetone was accomplished.
Column 2. 3% Poly-m-phenoxylene on 80/100 Chrom P DMCS, 6' x 1/8".
Injector 250°C, manifold 250°C, oven 130° - 200°C @ 8°/min.
FID, N- at 20 ml/min. Results: moderate resolution of
sample, 18 peaks, from raw oil.
Column 3. 10% Dow Corning High Vacuum Grease on 80/100 AWFB-DMCS
10' x 1/8". Injector 340°C, oven 150° - 350°C @ 10°/min
FID, N2 20 ml/min. Results: 48 peaks minimum resolution
from raw oil.
Column 4. 1% Polyphenylether (6 rings) on 80/100 AWFB-DMCS 3' x 1/8".
Injector 250°C, manifold 250°C, oven 130°C @ 10°/min FID,
N£ 20 ml/min. Results: moderate resolution of sample,
23 peaks from raw oil.
Column 5. 10% SP-2100 on 80/100 Suppelcoport 6' x 1/8". Injector
250°C, manifold 250°C, oven 60° - 250°C @ 5%nin. FID,
N2 20 ml/min. Results: Better resolution of components;
58 - 52 peaks from raw oil with better baseline separations.
Column 6. 10% Carbowax 20 M on 80/100 Supelcoport 6' x 1/8".
Injector 250°C, manifold 250°C, oven 60°- 250°C @ 5%nin,
FID, N£ 20 ml/min. Results: Good" resolution of low boil-
ing compounds.
DISTILLATION OF PYROLYTIC OILS
Distillation offers a possible method for processing and refining
pyrolytic oils obtained from lignocellulosic materials to yield more desir-
able and useful products of greater value, and thereby, increasing the
economic value of these oils. The oils contain a wide spectrum of organic
compounds including a large number of aromatic compounds. Because of the
wide variety of organic compounds in the oils, they offer the potential as
a source of chemical materials which should find many industrial applica-
tions .
30
-------
A number of distillation experiments were conducted with oils obtained
from the Tech-Air Corporation. These include distillation at atmospheric
pressure and at 0.2-0.4 mm mercury, fractional distillation at reduced
pressure, steam distillation and vacuum stripping. The data from these
experiments have been reported [12]. Representative liquid chromatograms
are presented in Figures 17, 18, 19, 20, 21, 22, and 23.
HYDROGENATION
Oil samples from different sources were hydrogenated catalytically to
determine how much hydrogenation would occur and the effect of hydrogenation
on the stability of the oil and to prepare samples for use in various separa-
tion schemes. Hydrogenation was carried out in a Parr Model 3911 hydrogena-
tion apparatus which provides for agitation by shaking and can be used at
pressures up to approximately 4 atmospheres. One hydrogenation was conducted
at atmospheric pressure utilizing a recycling of the hydrogen in a stirred
flask containing the sample and catalyst. Anhydrous ethanol was used as a
solvent, and five percent palladium on activated carbon or five percent plat-
inum on activated carbon was used as a catalyst. The results from the hydro-
genations with the low pressure Parr apparatus and at atmospheric pressure
are given in TABLE 6.
The data from hydrogenations 5,6 and 7 show that the Pd catalyst per-
forms better as the hydrogen absorbed is approximately fifty percent greater
in one-third of the time used for the hydrogenations with Pt. The data from
hydrogenation 4 show that hydrogenation at atmospheric pressure is too slow.
Examination of the data from hydrogenations 5, 8 and 9 shows that the Blue
IV fan oil from both hardwood and pine chips absorbed approximately the same
amount of hydrogen under similar conditions, whereas the Blue IV composite
hardwood oil adsorbed 2.2 times as much hydrogen as the Blue IV composite
pine oil. It is of interest that the vacuum stripped hardwood oil, hydro-
genation 11, absorbed 1.56 as much hydrogen as the vacuum stripped pine oil,
hydrogenation 10.
Hydrogenations are frequently carried out at a much higher pressure
than those discussed above. In order to test a higher initial hydrogen
pressure, a Parr Model 1108 calorimeter bomb was connected to a high pressure
hydrogen reservoir (lecture bottle size) utilizing a Parr oxygen bomb filter
hose assembly and stainless steel tubing. Agitation was provided by means
of a magnetic stirrer. Three hydrogenations were carried out with this
apparatus with vacuum stripped Blue IV fan pine oil. In each hydrogenation,
two grams of five percent palladium on activated carbon and 100 ml of abso-
lute ethanol were used. The hydrogenated oil was recovered by removal of
the catalyst by filtration and then vacuum stripping of the ethanol at 2 mm
pressure. The results of these three hydrogenations are given in TABLE 7.
An examination of the data shows that the hydrogen absorption is the same
for each experiment and that the samples absorbed approximately seventeen
percent more hydrogen than the same sample at approximately 4 atmospheres
(hydrogenation 10 TABLE 6).
31
-------
20
Minutes
35
Figure 17. Survey liquid chromatogram of combined fractions
from vacuum distillation.
10
15
Minutes
25
30
35
Figure 18. Survey liquid chromatogram of spinning band fraction one.
32
-------
0 5 10 15 20 25 30 35
Minutes
Figure 19. Survey liquid chromatogram of spinning band fraction five.
0 5 10 15 20 25 30 35
Minutes "
Figure 20. Survey liquid chromatogram of spinning band fraction nine.
33
-------
15 20
Minutes
25
30
35
Figure 21. Survey liquid chromatogram of condenser oil vacuum
stripped without heat.
10
15 20
Minutes
25
30
35
Figure 22. Survey liquid chromatogram of 100°-105°C organic layer
from steam distillation.
34
-------
Figure 23. Survey liquid chromatogram of 100 -105 C aqueous phase
from steam distillation.
35
-------
TABLE 6. HYDROGENATIONS AT MODERATE PRESSURE*
No. Sample Source
1 Coredle Condenser
Oil
2 Blue IV Hardwood
Composite Oil
3 Blue IV Pine
Composite Oil
4 Blue IV Fan
Hardwood Oil?
5 Blue IV; Fan
Hardwood Oil
6 Blue IV Fan
Hardwood Oil
7 Blue IVi-Fan
Hardwood Oil
8 Blue IV Fan
Pine Oil
9 Blue IV, Fan "-
Pine Oil
10 Blue IV Fan
Pine Oil, Vacuum
Stripped
11 Blue IV Fan
Hardwood Oil,
Vacuum Stripped
Weight
8
32.2
20.8
24.1
65.1
54.0
52.1
68.8
45.9
56.9
34.4
59.0
Water
%
19.5
' 12.8
17.8
12.4
12.4
,«-s
12.4
12.4
17.9
17.9
0
0
Weight
"dry"
oil?
26.0
18.1
19.8
57.0
47.3
45.6
60.3
37.7
46.7
34.4
59.0
Initial
pressure,
psig
55.2
55.5
56.0
Ambient
pressure
55.1
. f
55.3
56.2
57.2
57.1
58.1
59.0
Time
hrs
18
20
26
60
22
72
72
24
24
26
60
H2 Absorbed
mg/g on
"dry" basis
1.4
4.9
2.2
1.1
2.7
1.9
1.4
2.4
2.4
1.8
2.8
* 5% Pd on activated carbon was used in all experiments except 6 and 7, in
which 5% Pt on activated carbon was used. Two grams of catalyst were used
in each experiment. Approximately 200 ml of absolute ethanol was used for
each hydrogenation.
t Calculated dry weight of oil based on percent water.
^ This experiment was conducted in the recycle apparatus at ambient pressure.
36
-------
TABLE 7. HYDROGENATIONS AT INTERMEDIATE PRESSURE
Initial Pressure H£ Absorbed
No. Sample Source Atmospheres mg/g
12 Blue IV Fan Pine Oil 18.0 2.1
13 Blue IV Fan Pine Oil 19.5 2.1
14 Blue IV Fan Pine Oil 20.0 2.1
37
-------
EXPERIMENTAL—PHASE II
SEPARATION EXPERIMENTS
The objective of this phase on separation work with pyrolytic oils was
to obtain preliminary data on some approaches that could possibly be used for
development of a process that would produce more refined fractions of oil
that contain predominantly one chemical class of compounds. The broad classes
of chemical substances in raw pyrolysis oil are phenolics, aromatic neutral
compounds (neutrals of high aromaticity, NHA) , acidic compounds, and a group
of substances with "sugar-type" characteristics which are termed polyhydroxy
neutral compounds (PNC). The emphasis in the separation experiments has been,
therefore, to focus on obtaining fractions of the oil that contain essentially
one of the general classes of substances in the oils. This is a report of
the laboratory work of this phase at the bench level on a batch basis.
The five major approaches involving extraction techniques that were
tested are:
A - Extraction of oil sequentially with water at 25°C, 50°C, and 95°C.
B - Extraction of oil with sodium sulfate solution (salting-out effect).
C - Extraction of oil simultaneously with an organic solvent and water
(three phase system).
D - Extraction of sodium hydroxide soluble fractions of pyrolysis oil.
E - Extraction of organic solvent solutions of pyrolysis oil with
water.
Vacuum Stripping of Raw Oil
Based on a number of extraction and separation experiments on a batch
basis with raw and vacuum stripped pyrolysis oils, vaouum stripped oil gave
better results than the raw oils. The vacuum stripping provides for the
removal of the volatile organics and most of the water in the oil with poten-
tial subsequent recovery of these organic compounds. Our analysis show that
the major organic component in the volatile fraction is acetic acid. For
these reasons, our preliminary separation techniques are based on using vacuum
stripped oil. Figure 24 shows schematically the vacuum stripping of the oil
with yields.
38
-------
Crude Pyrolysis Oil, 100 g
1
Vacuum stripped at
2 mm and ambient
temperature
Vacuum Stripped Oil, 82.1 g Volatile Fraction
Water 10.8 g
Acids 7.1 g
Figure 24. Removal of volatiles from pyrolytic oil.
Extraction of Oil Sequentially with Water at 25°C, 50°C, and 95°C
A sample of vacuum stripped oil was extracted sequentially with water
at 25°C, 50°C and 95°C in an effort to separate the more water soluble sub-
stances. Figure 25 shows schematically this separation process and the
recovery of the different fractions are given in TABLE 8. The overall
recovery was good. The liquid chromatogram, Figure 26, shows that the water
extract is essentially free of the components of the oil which emerge in the
latter two-thirds of the liquid chromatogram of the raw oil, Figure 15. The
liquid chromatogram, Figure 27, shows that most of the components that appear
in the initial part of the liquid chromatogram of the raw oil has been
extracted sequentially with water at 25°C, 50°C, and 95°C. The liquid chro-
matograms of the water extract fractions at 50°C and 95°C were very similar
to Figure 26 of the 25°C water extract.
The significance of these results is that the oil can be separated
into water soluble and water insoluble fractions which offer the opportunity
for recovery of useful fractions of aromatic compounds. The water insoluble
fractions, based on our analysis, are composed of phenolics and neutral aro-
matics. The separation of this fraction into a highly concentrated phenolic
fraction and highly concentrated fraction of aromatic neutral compounds could
probably be accomplished by either fractional distillation or extraction
with alkaline solution. The aqueous phases could be combined and subjected
to a separation of the components with an aqueous salt solution as described
below to yield a fraction with mainly phenolics and another fraction with
mainly polyhydroxy neutral substances.
Extraction of Oil with Sodium Sulfate Solution
An extraction experiment with a sodium sulfate solution (90% saturated)
was conducted to determine if extraction with aqueous salt solutions would
offer a useful separation of the oil. The schematic for this extraction is
shown in Figure 28, and the overall recovery was good.
39
-------
Vacuum Stripped Pyrolysis Oil, 82.1 g*
L
KM-u-n
Water at 25°C
Insoluble Organic
Fraction- 73.1%
Water at 50°C-
Insoluble Organic
Fraction- 63.8%
Water at 95°C-
Mixer and
Separation
Aqueous Fraction
Phenolics - 8.2%
Polyhydroxy neutrals-18.8%
Mixer and
Separation
Aqueous Fraction
Phenolics - 2.1%
Polyhydroxy neutrals-7.2%
Mixer and
Separation
Insoluble Organic
Fraction - 49.2g
Phenolics-12.2%
Aromatic neutrals-47.7%
Aqueous Fraction
Phenolics - 0.7%
Polyhydroxy neutrals-3.2%
82.Ig of vacuum stripped oil was obtained from lOOg of this raw oil.
Figure 25. Extraction of oil sequentially with water at 25°C, 50°C, and 95°C.
40
-------
Figure 26. Liquid chromatogram of 25 C water extract of pyrolytic oil.
Figure 27. Liquid chromatogram of pyrolytic oil after successive extraction
with water at 25°C, 50°C, and 95°C.
41
-------
TABLE 8. YIELDS OF FRACTIONS FROM WATER EXTRACTION OF OIL
Phenolics
Water Insoluble
Fraction
10 g
25°
6.7 g
Water Soluble
50°
1-7 g
Fractions
95°
0.6 g
Total
9.0 g
Aromatic
neutrals 39.2 g
Polyhydroxy
neutrals
Totals 49.2 g
—
15.4 g
22.1 g
—
5.9 g
7.6 g
—
2.6 g
3.2 g
—
23.9 g
32.9 g
Vacuum Stripped Pyrolysis Oil, 82.1 g
I
f
Insoluble Fraction
Phenolics- 12.7%
Aromatic and
polyhydroxy neutrals-71.5%
Saturated
SO. Solution "
4
Mixer and
Separation
1
Aqueous Fraction
Phenolics- 11.1%
Polyhydroxy neutrals-4.8%
Figure 28. Extraction of pyrolytic oil with sodium sulfate solution.
The importance of these results is that with the sodium sulfate solu-
tion approximately 82% of the polyhydroxy neutrals are in the insoluble
fraction with about 18% in the aqueous fraction. The phenolics are approxi-
mately 70% of the organics in this aqueous fraction. There are two approaches
that can be used involving the sodium sulfate extraction. One approach would
be to use the sodium sulfate extraction as the first step as shown in Figure
28 to produce an aqueous fraction of mainly phenolics. The insoluble
organic fraction would then be treated with water extraction as depicted in
Figure 25 to remove the polyhydroxy neutrals. The other approach would
be to treat the oil as outlined in Figure 25, and then the three aqueous
fractions would be combined followed by the addition of sodium sulfate. This
approach could possibly provide a good separation between the phenolics and
the polyhydroxy neutrals. The addition of a water insoluble organic sol-
vent may be necessary in such a step to serve as a solvent for the poly-
hydroxy neutrals.
42
-------
Extraction of Oil Simultaneously with Organic Solvents and Water: Three
Phase System
Organic solvents offer a good potential for effecting separation of
pyrolysis oils into fractions which contain very similar organic compounds.
Some extractions with diisopropyl ether and anisole (methylphenyl ether) were
tried with vacuum stripped oil. It was found difficult to have good contact
of the organic solvent with only the oil because of the increase in the
viscosity of the oil. Addition of an equal volume of water to the mixture
produced a nonviscous three-phase-system containing an ether phase, an
aqueous phase and a heavy oil phase with an overall recovery of approximately
96%. The schematic for diisopropyl ether and water separation along with
yields is shown in Figure 29 and the schematic for anisole and water, Figure
30. Based on our analysis, the phenolics in the water fraction are mainly
dihydroxy phenols; in the diisopropyl ether phase, alkylphenols; and in oil
phase, ether phenols. The aqueous phases from both of the diisopropyl
ether-water separations could be combined and possibly separated into a
highly concentrated phenolic fraction by salting out the polyhydroxy neutrals
with addition of sodium sulfate or some other salt.
In the anisole experiment, the phenolics were evenly divided between
the anisole fraction and the aqueous fraction with a small amount in an oil
insoluble fraction. About 88% of the aromatic neutrals were extracted into
the anisole fraction, which contained about 62% of the original charge. A
good potential step for processing this fraction would be fractional distil-
lation. The oil insoluble fraction, which contained about 8.4% of the origi-
nal charge, was approximately 85% aromatic neutrals and could be further
processed by fractional distillation. The aqueous phase could be treated by
the salting out technique with sodium sulfate as shown in Figure 27 to yield
a highly concentrated phenolic fraction.
Extraction of Sodium Hydroxide Soluble Fractions of Pyrolysis Oil
A sample of vacuum stripped pyrolysis oil (154 g) was treated with 300
mi of 2% sodium hydroxide solution and approximately 52.6% dissolved. A
series of methylene chloride extractions then were made at three different pH
ranges. The "insoluble oil phase" upon treatment with additional 2% sodium
hydroxide solution, dissolved in 400 ml of the alkaline solution. This solu-
tion was subjected to a series of methylene chloride extractions at the same
pH ranges. The schematic for these extractions were presented in Figure 31.
! The overall recoveries were good, and the yield data are presented in TABLE
9. An examination of the data shows that phenolics are obtained with methyl-
ene chloride at each pH range and approximately 52% of the phenolics remain
in the aqueous phase at pH range 1 to 3. The significance of this experiment
is that the pyrolysis oil will dissolve in sufficient sodium hydroxide solu-
tion which offers the opportunity for a series of extractions at different pH
ranges and also with a variety of organic solvents.
43
-------
Vacuum Stripped Pyrolysis Oil, 100 g
100 ml Water-
100 ml Diiso-
propyl ether
Mixer and
Separation
Oil Fraction
Phenolic - 6%
Aromatic
neutrals-24.9%
Ether Fraction
Phenolic - 5%
Aromatic
neutrals-16.7%
I
Aqueous Fraction
Phenolic - 11.2%
Polyhydroxy neutrals-34.3%
Diisopropyl
ether
Mixer and
Separation
Oil Fraction
Ether Fraction
Phenolic- 1.5%
Water
Mixer and
Separation
I -
Oil Fraction
Phenol!cs-4.4%
Aromatic
neutrals-16.2%
Aqueous Fraction
Phenolic - 1.5%
Neutrals* - 5.2%
Chemical nature unknown.
Figure 29. Combined diisopropyl and water extraction of pyrolytic oil.
44
-------
Vacuum Stripped Pyrolysis Oil, 110 g
100 ml Water —
100 ml Anisole
Three Phases
Mixer and
Separation
Anisole Fraction*
Phenolic - 12.5 g
Aromatic neutrals - 56 g
Aqueous Fraction
Phenolic - 12.5 g
Polyhydroxy neutrals-32.1 g
Oil Insoluble Fraction
Phenolic - 1.4 g
Aromatic neutrals - 7.8 g
The removal of all anisole from this fraction was difficult so that total
recovery is greater than 100%.
Figure 30. Combined anisole and water extraction of pyrolytic oil.
Extraction of Organic Solvent Solutions of Oil
The vacuum stripped pyrolysis oil dissolves in methylene chloride and
in n-butanol to give complete solutions. Solutions of vacuum stripped oil in
methylene chloride were extracted with water and the combined water extracts
were then extracted in one experiment with diisopropyl ether and in a second
experiment with methyl isobutyl ketone (MIBK). The schematics for these two
experiments are shown in Figures 32 and 33. The data are summarized in
TABLE 10. The significance of the data in these experiments is that a
fraction of phenolics is obtained with MIBK which contains less than 10%
other organics. An examination of the data will also indicate one of the
difficulties encountered in working with pyrolysis oils. One would expect
the quantity of phenolics in the final methylene chloride fractions to be in
closer agreement. The lack of agreement can be attributed to differences in
experimental techniques and to the need of improvement in analytical tech-
niques .
45
-------
Vacuum Stripped Pyrolysis Oil, 154 g
300 ml 2% NaOH-
Solution
1
Mixer and
Separation
Aqueous Solution
pH 8 to 10
Organics- 81 g (52.6%)
Insoluble Oil Phase
Organics- 74 g (49.1%)
(see next page)
CH2C12-
Aqueous Solution
Adjust pH 5 to 7
CH2C12-
Aqueous Solution
Adjust pH 1 to 3
Mixer and
Separation
Mixer and
Separation
Mixer and
Separation
Aqueous Solution
Phenolics - 9.9%
Polyhydroxy neutrals-70.1%
CH2C1_ Extract
Phenolic- 1.8%*
Aromatic neutrals - 7.!
CH2C1 Extract
Phenolic- 1.1%
Aromatic neutrals -2.0%
CH-C1 Extract
Phenolic -5.3%
Aromatic neutrals - 0.7%
Percent yield is based on weight of material extracted from 81 g of organics.
Figure 31. Extraction of pyrolytic oil with 2% sodium hydroxide solution.
46
-------
Vacuum Stripped Pyrolysis Oil, 154 g (cont'd)
Insoluble Oil Phase, 74 g
400 ml 2% NaOH.
Solution
Aqueous Solution
Adjust pH 5 to 7
Aqueous Solution
Mixer and
Separation
Aqueous Solution
Phenolic - 12.5%
Polyhydroxy neutrals-9.2%
i
Aqueous Solution
pH 8 to 10
Organics - 74 g
Mixer and
Separation
Mixer and
Separation
Mixer and
Separation
CH Cl Extract
Phenolic- 2.5%*
Aromatic neutrals - 28.3%
CH Cl Extract
Phenolic - 5.1%
Aromatic neutrals -8.9%
CH Cl Extract
Phenolic -5.3%
Aromatic neutrals -2.4%
Percent yield is based on weight of material extracted from 74 g organics.
Figure 31 (cont'd). Extraction of pyrolytic oil with 2% sodium hydroxide
solution.
47
-------
TABLE 9. YIELDS FROM METHYLENE CHLORIDE EXTRACTIONS OF ALKALINE SOLUTIONS
OF PYROLYTIC OIL
First Series Second Series
Fraction
Extractions
Weighting
Extractions
Weighting
Total
Yield
%
Yield
pH 8 to 10
Phenolics
Aromatic neutrals
pH 5 to 7
Phenolics
Aromatic neutrals
pH 1 to 3
Phenolics
Aromatic neutrals
Aqueous Phase
Phenolics
Polyhydroxy neutrals
Tar neutrals
1.46
6.4
0.89
1.62
4.29
0.57
8.02
56.8
1.85
20.9
3.77
6.56
3-92
1.78
9.25
6.81
17.5
3.31
23.3
4.67
8.21
8.21
2.35
17.3
63.6
17.5
2.17
17.9
3.06
5.38
5.38
1.54
11.3
41.7
11.8
Totals
Phenolics
Aromatic neutrals
Polyhydroxy neutrals
Tar neutrals
33.5
37.9
63.9
17.5
21.9
24.8
41.7
11.8
The vacuum stripped pyrolysis oil is soluble in n-butanol, and an
aqueous extraction experiment with a n-butanol solution of pyrolysis oil was
carried out to determine the distribution of the phenolic and other organics
between the aqueous and n-butanol fractions. The schematic for this experi-
ment with yields for each fraction is given in Figure 34. The important
result of this experiment is the reduced amount of polyhydroxy neutrals
in the aqueous phase as compared with the other extractions with the excep-
tion of the sodium sulfate extraction. There is the potential that extraction
of a n-butanol solution of pyrolysis oil with sodium sulfate solution could
yield an aqueous solution with a high concentration of phenolics relative to
other organics. In this experiment, material recovery is not too good because
in the removal of the n-butanol at low vacuum, some of the more volatile aro-
matic compounds were lost.
48
-------
Vacuum Stripped Pyrolysis Oil, 101 g
Methylene
Chloride
Dissolve in
Methylene Chloride
Water-
Mixer and
Separation
I
Aqueous Fraction
Phenolics -11.5%
Polyhydroxy neutrals-33.3%
Methylene Chloride Fraction
Phenolics - 13.4%
Aromatic neutrals - 41.3%
Diisopropyl
ether
Mixer and
Separation
Ether Fraction
Phenolics- 2.2%
Aqueous Fraction
Phenolics - 12.3%
Polyhydroxy neutrals-28.1%
Figure 32. Extraction of methylene chloride solution of pyrolytic
oil with water followed by diisopropyl ether extraction of
aqueous frac t ion.
49
-------
Vacuum Stripped Pyrolysis Oil, 100 g
Methylene_
Chloride
Dissolve in
Methylene Chloride
Water•
Mixer and
Separation
f
Aqueous Fraction
Phenolic- 16.3%
Polyhydroxy neutrals-32.4%
Methylene Chloride Fraction
Phenolics - 8.8% /
Aromatic neutrals - 42.9%
Methyl isobutyl
ketone
Mixer and
Separation
Ketone Fraction
Phenolics-8.8%
Aromatic neutrals - 0.8%
Aqueous Fraction
Phenolics - 7.5%
Polyhydroxy neutrals-30.1%
Figure 33. Extraction of methylene chloride solution of pyrolytic oil
with water followed by methylisobutyl ketone extraction
of aqueous fraction.
50
-------
TABLE 10. YIELDS IN FINAL FRACTIONS FROM SEPARATION TECHNIQUES
IN FIGURES 32 AND 33
Final Fraction
Methylene chloride
Phenolics
Aromatic neutrals
Aqueous
Phenolics
Polyhydroxy neutrals
Organic solvent
Phenolics
Aromatic neutrals
Diisopropyl Ether
Experiment
13.5 g
41.7 g
12.4 g
28.4 g
2.2 g
0
Methylisobutyl
Ketone Experiment
8.8 g
42.9 g
7.5 g
30.1 g
8.8 g
0.8 g
Vacuum Stripped Pyrolysis Oil, 100 g
n-Butanol•
Water
Dissolve in
n-butanol
Mixer and
Separation
Aqueous Fraction
Phenolics- 10.7 g
Polyhydroxy neutrals-13.6 g
n-Butanol Fraction
Phenolics - 11.2 g
Aromatic neutrals and
Polyhydroxy neutrals-49 g
Figure 34. Extraction of n-butanol solution of pyrolytic oil with
water.
51
-------
EXPERIMENTAL—PHASE III
PYROLYTIC OIL
The pyrolytic oil for this experimental phase was taken from the oil
produced in a run in the Georgia Tech pyrolysis pilot plant (capacity, 225
kg/hr) on October 12, 1978. The converter feedstock was pine chips dried to
contain approximately six percent moisture, and the air-to-feed input ratio
was continually adjusted within a narrow range to maintain a temperature of
125° to 130°C in the off-gases passing from the headspace of the reactor to
the condensers. The condenser temperatures were held near 75°C. These
closely controlled low temperatures resulted in less thermal cracking than
had been observed in earlier converter runs with higher temperatures. "
The selected containers of pyrolytic oil were stirred thoroughly and
the moisture content of the oil in each container was determined. Two four-
liter reference samples were taken from each container and stored in tightly
capped plastic containers for future reference. The remaining oil was com-
bined and thoroughly mixed. Eight four-liter samples were stored in tightly
closed plastic containers for laboratory work. The remaining oil was stored
in tightly closed plastic lined containers as a reserve supply.
Characterization of Pyrolytic Oil Sample
The percent moisture in the sample was determined by azeotropic dis-
tillation with toluene (Dean and Stark Method). The percent solid material,
mainly fine fiber and char fines, was determined by dissolving a weighed por-
tion of the oil in a large excess of acetone and passing the solution through
a tared glass filter paper. The filter paper and residue were thoroughly
washed with acetone, dried, and weighed. The percent ash was determined by
charring weighed oil samples in tared crucibles by means of an infra-red
lamp, igniting the char in a muffle furnace, and determining the weight of
the ash. Sulfur was determined by igniting two-gram oil samples at 30 "
atmospheres in an oxygen bomb calorimeter. No turbidity was observed when
barium chloride was added to filtered washings from the oxygen bomb, and no
increase was observed in the weight of tared Gooch crucibles used to filter
the solution of barium chloride in the washings. The density of the mixed
oil sample was calculated from the weight of 200 ml at 25°C. The percent of
carbon, hydrogen and nitrogen was determined using a Perkin Elmer Model 240
Elemental Analyzer. Results of these characterizations are shown in TABLE 11.
SEPARATION EXPERIMENTS
The results of the experimental work in Phase II with different extrac-
tion techniques with pyrolytic oil were carefully evaluated for further
investigation for the development of a pilot plant concept for processing
52
-------
TABLE 11. PROPERTIES OF PYROLYTIC OIL SAMPLE
Determined
Percent Moisture
Percent Solids
Percent Ash
Percent Sulfur
Percent Carbon
Percent Hydrogen
Percent Nitrogen
Density (g/m£)
— — '•""• • - . ... — - — -
Sample 1
14.7
0.38
0.055
<0.001
57.27
6.72
0.06
1.234
Sample 2
14.9
0.43
0.054
<0.001
57.34
6.76
0.06
1.234
-- - • • —
Average
14.8
0.41
0.055
<0.001
57.30
6.74
0.06
1.234
pyrolytic oils. The selected processes were aqueous extraction (Process No.
1) , simultaneous extraction with water and an organic solvent (Process No. 2),
and dissolution of the pyrolytic oil in an organic solvent followed by aque-
ous extraction of the solution (Process No. 3). The first efforts were with
batch experiments of all three processes in which both vacuum stripped and
unstripped oil samples were examined and the effects of both polar and non-
polar solvents were studied. Based on the results of the batch experiments,
Process No. 1 and Process No. 2 using a polar organic solvent were chosen
for:continuous countercurrent extractions of both vacuum stripped and
unstripped pyrolytic oil. The batch experiments will be described first
followed by the description of the continuous extraction experiments.
Initial Batch Separation Procedures
The batch separations were performed by stirring approximately 100,
200, or 500 g of oil, weighed to the nearest 0.1 g, with the extracting sol-
vent system for 30 minutes in a tall form 1,000 ml beaker at approximately
900 revolutions per minute using a 4 cm PTFE coated bar with a magnetic
stirrer. At the end of the contact period the beaker was chilled to immobil-
ize the insoluble tar phase so that the extracting solvent phase or phases
could be removed by decantation. Conventional separately funnels were used
to separate the aqueous and immiscible organic solvent phases.
Process No. 1. Water Extraction Procedures—
Six samples of pyrolytic oil were extracted with water, and the water
phases were separated from the insoluble organic phases by decantation. Two
additional aqueous extractions were made, each using the insoluble organic
phase from the preceding extraction. A schematic flow diagram of this pro-
cedure is shown in Figure 35, x^hich shows the treatment of unstripped pyro-
lytic oil by vacuum stripping and subsequent water extraction as solid lines
at the top of the figure and by water extraction without vacuum stripping as
a broken line at the top of the figure. The broken lines at the bottom of
the figure indicate generalized further treatments of the separated phases.
Six samples of oil were extracted with water as listed below. The
aqueous phase and insoluble organic phase from Extraction I (1) were used to
53
-------
UNSTRIPPED PYROLYTIC OIL'
•Vacuum Strip1
L
STRIPPED
OIL
VOLATILES
PHASE
Water
extraction
INSOLUBLE PHASE 1
AQUEOU
S PHASE 1 —i
Water extraction
INSOLUBLE PHASE 2
Water extraction
AQUEOUS PHASE 2-
INSOLUBLE PHASE 3
AQUEOUS PHASE 3-
TO ANALYSIS
Extractions,
distillations,
etc.
Combine
r
Extract weighed
portion with^
3 portions of
organic solvent
FRACTIONS
COMBINED
ORGANIC
FRACTION
i
COMBINED
AQUEOUS
PHASE
ANALYTICAL
SAMPLE
COMBINED
AQUEOUS
FRACTION
Samples and yields shown in UPPER CASE LETTERS
'Operations shown in lower case letters
Figure 35. Aqueous batch extraction, Process No. 1.
develop analytical techniques at Georgia Tech and at Atlanta University. The
fractions from Extraction I (2-6) were used to experiment with techniques to
obtain additional fractions. The separation techniques are described in a
later section of this report.
-------
0 Extraction I (1)—A 102.9 g sample of vacuum stripped oil was
extracted with three 100 ml portions of deionized water.
o Extraction I (2)—This experiment was a duplicate of I (1) to pro-
vide a water solution for subsequent extraction with a polar organic
• solvent.
o Extraction I (3)—This experiment was run as I (1) and I (2) to pro-
vide an aqueous solution for extraction with a nonpolar organic
solvent.
0 Extraction I (4)—This extraction was performed as I (1) except that
a 203.4 g portion of unstripped oil was extracted with two 200 ml
portions of water. The water extract was reserved for contact exper-
iments with activated carbon.
0 Extraction I (5)—This experiment was similar to I (4) .
o Extraction I (6)—A 400 g unstripped oil sample was extracted with
400 ml water followed by two successive extractions with 200 ml
^ portions of water.
The "water solution fraction and water insoluble fraction were used for further
analysis and testing of additional separation techniques.
No attempt was made to isolate individual compounds from the large
number present in each separated phase or fraction. Quantitative analysis
was directed only toward separating and identifying classes of compounds
having similar solubilities or measurable chemical properties, which might
be related to their potential commercial use. Based on analytical methods,
which will be described in a later section of this report, the vacuum strip-
ping and extraction yields were determined as volatile organics, nonvolatile
organics (NVO), phenolics, polyhydroxy neutral compounds (PNC) and neutrals
of high aromaticity (NBA). The polyhydroxy neutral compounds were estimated
by subtracting the phenolics in the water phases or fractions from the
corresponding total nonvolatile organics. Neutrals of high aromaticity were
estimated by subtracting the phenolics from the total organics in an organic
solvent phase or fraction. The results of the batch extraction, expressed
as percent of the moisture-free unstripped oil sample, are shown in TABLE 12.
Since the moisture free oil contained seven percent volatile compounds the
total nonvolatile organics should approach 93 percent.
The percent nonvolatile organics (NVO) was determined by removing the
solvent from a weighed sample of the separated phase on a rotary vacuum
evaporator with caution to avoid heating. It is believed that incomplete
solvent removal from the organic phase led to the apparently high total NVO
values in Extractions I (2) and I (6). The aqueous phase from Extraction
I (2) was extracted with three successive portions of methylisobutyl ketone
(MIBK). The MIBK extracts were combined to form the MIBK fraction. The
distributions of the classes of organic compounds in the MIBK fraction and
the extracted water fraction are shown in parentheses. The distributions
'resulting from a similar extraction of the water phase in Experiment I (3)
with-chloroform are represented in a similar manner. The percent NVO was
determined separately for- each of the four successive water phases in
Extraction I (6) to show the quantity of organic material removed by each
extraction step. Since most of the water soluble material was found in the
55
-------
TABLE 12. COMPOSITION OF YIELDS FROM BATCH WATER EXTRACTIONS, PROCESS NO. 1
Extraction Experiment
Extraction I (1)
Aqueous Phase
Insoluble Organic Phase
Extraction I (2)
Aqueous Phase
Aqueous Fraction
MIBK Fraction
Insoluble Organic Phase
Extraction I (3)
Aqueous Phase
Aqueous Fraction
Chloroform Fraction
Insoluble Organic Phase
Extraction I (4)
Aqueous Phase
Insoluble Organic Phase
Extraction I (5)
Aqueous Phase
Insoluble Organic Phase
Extraction I (6)
First Aqueous Phase, 1(6) Al
Second Aqueous Phase, I (6) A2
Third Aqueous Phase, I (6) A3 -
Fourth Aqueous Phase, I (6) 4
Insoluble Organic Phase, I (6)0
Percent
NVO*
53.8
39-8
50.4
(38.5)
(11.9)
55.5
52.8
(41.7)
(11.1)
39.9
40.1
51.4
52.3
41.4
44.6
6.0
2.5
0.1
55.8
Percent
Phenolic
28.7
13.7
34.0
(24.0)
(10.0)
5.5
41.4
(33.2)
(8.2)
23.3
Percent
PNCt
25.1
-
14.5
(14.5)
-
12.2
(12.2)
-
Not Determined [Stock
Not Determined [Stock
Not Determined [Stock
Not Determined [Stock
13.6
31.0
Percent
NHAT
-
26.1
1.9
(1.9)
50.0
2.9
(2.9)
15.6
K4)A]
1(4)0]
K5)A]
1(5)0]
-
Not Determined
Not Determined
Not Determined
11.9
-
43.9
* Non Volatile Organics
t Polyhydroxy Neutral Compounds
$ Neutrals of High Aromatic!ty
56
-------
first water phase, 1(6) Al, only this phase was analyzed and reserved for
further experiments.
Process No. 2. Three Phase Extraction Procedure—
In the three phase extraction technique the oil sample was extracted
with a vigorously stirred mixture of water and an immiscible solvent. The
liquid phases were decanted from the insoluble tar phase and separated into
aqueous and organic phases. A schematic diagram of this process is shown
in Figure 36.
UNSTRIPPED PYROLYTIC OIL* »~ Vacuum Strip1-
f t
| STRIPPED VOLATILES
OIL PHASE
Extract with
mixture of water
and organic solvent
J I
I J
WATER ORGANIC INSOLUBLE
PHASE PHASE TAR
Samples and yields shown in UPPER CASE LETTERS
t Operations shown in lower case letters
Figure 36. Three phase extraction, Process No. 2.
Four batch extractions were performed using mixtures of water with
MIBK as a polar organic solvent or water with chloroform as a nonpolar sol-
vent as follows.
0 Extraction II (1)—A 103.1 g sample of vacuum stripped oil was
stirred with a mixture of 1QO mT water and 100 ml MIBK. The
mixture was allowed to stand, and the water and organic phases
were separated.
0 Extraction II (2)—Extraction II (2) was performed as II (1) using
105.6 g unstripped oil, 200 ml chloroform, and 100 ml water.
o Extraction II (3)—This extraction was similar to II (1) except
that the sample was 97.9 g unstripped oil.
o Extraction II (4)—This experiment was run in the same manner as
II (1).
57
-------
The distributions of the main classes of compounds were determined
following the scheme described above for Process No. 1. These distributions
are shown in TABLE 13. The letter codes, e.g., II(1)A, shown after each
phase are included to facilitate their identification as starting materials
for additional experiments to be described in later sections of this report.
Process No. 3. Dissolution in an Organic Solvent Followed by Water
Extraction—
In these experiments listed below, the oil sample was dissolved in an
organic solvent, and the resulting solution was extracted with water. A
schematic diagram of this process is shown in Figure 37.
o Extraction III (1)—A 102.6 g sample of vacuum stripped oil was
stirred with 200 ml chloroform. The chloroform solution was
extracted with three 100 ml portions of water.
° Extraction III (2)—This experiment was similar to III (1) except
that 200 ml MIBK was used to dissolve the oil, and the three water
extractions were carried out with a weighed fraction of the MIBK
solution with proportionally smaller quantities of water.
° Extraction III (3)—This experiment was similar to III (2) except
that the sample was unstripped oil.
The distributions of the identifiable classes of compounds were deter-
mined. These distributions are shown in TABLE 14. The chloroform insoluble
material in Extraction III (1) was readily soluble in acetone or five percent
•aqueous alkali, which indicates that the neutral material in the insoluble
tar phase contained a large number of hydroxyl groups. This interpretation
was supported by infra-red examination. The MIBK insoluble tars in Extrac-
tions III (2) and III (3) were readily soluble in acetone but only partially
dissolved in five percent aqueous alkali. With the support of infra-red
evidence it was concluded that these MIBK insoluble tars were a mixture of
polyhydroxy compounds and neutrals of high aromaticity.
Cbntinuous Countercurrent Extraction Procedures
Results of the batch extraction experiments indicated that continuous
countercurrent extraction work should be concentrated on Process No. 1 with
subsequent extraction of the resulting solution phase with MIBK and on
Process No. 2 using water and MIBK. Four experimental runs were made using
Process No. 1 with vacuum stripped oil, Process No. 1 with unstripped oil,
Process No. 2 with vacuum stripped oil and Process No. 2 with unstripped oil.
Modular construction was chosen for the countercurrent extractor to
permit relocation of the inlet and outlet points and to permit variations in
the length of the unstirred phase separation zones. A schematic diagram of
the counter current extractor is shown in Figure 38.
The apparatus consisted of a vertical tube with a heavy tar outlet at
the bottom and side inlets for solvent admission and a recycling line out-
let in the lower sections of the tube. The diameter of the mixing chamber
ras larger than that of the settling zones to prolong the residence time of
58
-------
_TABLE 13. COMPOSITION OF YIELDS FROM BATCH THREE PHASE EXTRACTIONS,
PROCESS NO. 2
Extraction Experiment
Extraction II (1)
Aqueous Phase 11(1) A
MIBK Phase II(1)M
Insoluble Tar Phase
Extraction 11(2)
Aqueous Phase II (2) A
.Chloroform Phase II(2)C
Insoluble Tar Phase
Extraction 11(3)
Aqueous Phase II (3) A
MIBK Phase II(3)M
Insoluble Tar Phase
Extraction 11(4)
Aqueous Phase II (4) A
MIBK Phase II(4)M
Insoluble Tar Phase
Percent
NVO*
39.2
53.0
0.1
41.0
51.7
2.2
38.7
50.7
2.1
41.6
54.1
0.5
Percent Percent
Phenolic PNCt
16.9 22.3
22.1
ND*
17.9 23.1
21.3
ND
ND
ND
7.9 33.7
27.6
ND
Percent
NBA**
-
30.9
-
30.4
-
26.5
Non volatile hydrocarbons
"fpolyhydroxy neutral compounds
**Neutrals of high aromaticity
determined
59
-------
UNSTRIPPED PYROLYTIC OIL
•*• Vacuum Strip
I
STRIPPED
OIL
VOLATILES
PHASE
Dissolve in
organic solvent
ORGANIC SOLUTION
PHASE
INSOLUBLE TAR
Extract
three times
with water
EXTRACTED
ORGANIC FRACTION
COMBINED
WATER FRACTION
Samples and yields shown in UPPER CASE LETTERS.
Operations shown in lower case letters.
Figure 37. Sequential organic water extraction, Process No. 3,
60
-------
Extraction Experiments
Extraction III(l)
Chloroform Phase III(1)C
First Aqueous Fraction III(1)A1
Second Aqueous Fraction III(1)A2
Third Aqueous Fraction III (1) A3
Extracted Chloroform Fraction III(1)CE
Insoluble Tar Phase III(1)MR
Extraction 111(2)
MIBK Phase III(2)M
First Aqueous Fraction III(2)A1
Second Aqueous Fraction III(2)A2
Third Aqueous Fraction III(2)A3
Extracted MIBK Fraction III(2)ME
Insoluble Tar Phase III(2)MR
Extraction 111(3)
First Aqueous Fraction III(3)A1
Second Aqueous Fraction III(3)A2
Third Aqueous Fraction III (3) A3
Extracted MIBK Fraction III(3)ME
Insoluble Tar Phase III(3)MR
' — «
Percent
NVO*
85.7
(27.5)
(3.5)
(1.4)
(36.6)
7.8
77.1
(23.1)
(4.9)
(4.9)
(44.2)
15.3
AS 1
D D • J_
(17.9)
(3.9)
(1.6)
(41.1)
37.9
Percent
Phenolics
20.3
(12.8)
(2.5)
(0.8)
(6.4)
4.5
23.4
(12.0)
(4.9)
(0.5)
(6.5)
' 3.7
(9.7)
(3.9)
(1.4)
(7.6)
8.0
Percent
PNC+
(14.7)
(1.0)
(0.6)
3.3
(11.1)
(0)
(0)
(8.2)
(0)
(0.2)
Percent
NHA**
65.4
(30.2)
-
53.7
(37.7)
11.6
(33.5)
29.9
* Non volatile organic
t Polyhydroxy neutral compounds
**Neutrals of high aromaticity
-------
T
20 cm
22 cm
Recycle
63 cm
Circulating
Pump
10 cm od
^
oo
V>
4.5 cm od
Solution Out
Stagnant Zone
Oil In
Stirring Zone
Settling Zone
Water In
MIBK In
Insoluble Oil Out
Figure 38. Countercurrent extractor.
62
-------
the oil and solvent in the vigorously stirred zone. A recycling line was
provided to withdraw a portion of the stream containing undissolved oil
droplets and return it to the top of the mixing chamber. The solvent supply
rate and the recycle flow rate were controlled by means of "Masterflex"*
variable speed tubing pumps.
The oil sample was led through the top of the stirring chamber to a
point level with the blades of a high speed propeller type stirrer. The
undissolved oil droplets settled downward through the tube countercurrent to
the incoming solvent stream. A portion of the rising solution phase and
descending undissolved oil droplets was withdrawn through the recycle loop
at a flow rate fifty times greater than the solvent intake rate and returned
to the top of the stirring zone. The heavy extracted oil phase was collected
and discharged at the bottom of the apparatus. The oil solution passed
through a constriction at the top of the mixing chamber into the stagnant
zone, and the dissolved oil stream flowed from an outlet near the top of the
tube.
Process 1A. Continuous Countercurrent Water Extraction of Unstripped Oil —
The experimental procedure was as follows . The apparatus was filled
with deionized water, and the stirrer and pumps were turned on. Unstripped
pyrolytic oil was admitted to the apparatus. The undissolved oil was with-
drawn from the bottom of the extractor and portions of the solution phase
which eluted from the top of the extractor were analyzed for dissolved non-
volatile organics (NVO) . The system was considered to be equilibrated when
no change was observed in the NVO concentration of the successive portions
of the diluted aqueous phase.
.1
Pyrolytic oil and water were fed into the equilibrated reactor at
carefully controlled rates from calibrated reservoirs. The undissolved oil
phase and the diluted water solution were collected in tared receiving
vessels and weighed, and the percent NVO in the oil and aqueous phases was
determined. The inputs and yields from this three hour experiment are shown
in TABLE 15.
The apparent loss of nonvolatile organic material was believed to be
distributed between the adherent tar and the solution remaining in the
extractor. The loss in water and volatile organics was attributed to evapo-
ration and leakage.
Process IB. Continuous Countercurrent Water Extraction of Vacuum Stripped
This experiment was conducted by the same method as Process LA using
vacuum stripped oil. The duration of the experiment was two hours and ten
minutes. The inputs and yields for this experiment are summarized in TABLE
16.
Cole Farmer Instrument Company, Chicago, Illinois.
63
-------
TABLE 15. INPUTS AND YIELDS. PROCESS 1A
Operation
Total g
g/minute
(average)
Inputs
Oil Sample In
Total Sample
Non Volatile Organic
Volatile Organic
Sample Moisture
Extraction Solvent In
Water
Outputs
Aqueous Phase
Non Volatile Organics
Water and Volatile Organics
Insoluble Oil Phase
Non Volatile Organics
Water and Volatile Organics
1,998
1,698
140
160
4,120
4,806
808
3,998
1,121
818
303
10.8
9.2
0.8
0.9
22.3
25.8
4.4
44.4
6.1
4.4
1.6
Apparent Losses
Total
Non Volatile Organics
Water and Volatile Organics
191.0
72.0
119.0
1.0
0.4
0.6
The loss of nonvolatile organics may be attributed to trapped tars in
the extractor and to dissolution in the remaining liquid phase. Evaporation
is believed to be the major cause of water and volatile losses.
Process 2 A. Continuous Countercurrent Three Phase Extraction of Unstripped
Oil—
The apparatus was filled by pumping in approximately equal parts by
volume of MIBK and water. The stirrer and recirculating pump were then
turned on, and the two solvent phases were thoroughly mixed for thirty min-
utes. The extractor was then equilibrated by passing in constant rate
streams of unstripped oil, MIBK, and water until 800 ml of oil had passed
into the extractor and the NVO concentration in the effluent stream was con-
stant. At this point the levels of oil, MIBK and water in the calibrated
feed reservoirs were recorded and the effluent stream was diverted into a
tared receiver. The levels of oil and solvents in the reservoirs and the
weight of the receiver were recorded at approximately five minute intervals
to insure constant input and output rates. At the end of 90 minutes the
final oil and solvent levels and the weight of the collected effluent were
64
-------
TABLE 16. INPUTS AND YIELDS, PROCESS IB
Operation
Total g
g/minute
(average)
Inputs
Oil Sample
Total Sample
Non Volatile Organic
Volatile Organic
Sample Moisture
Extraction Solvent In
Water
Outputs
Aqueous Phase
Non Volatile Organics
Water and Volatile Organics
Insoluble Oil Phase
Non Volatile Organics
Water and Volatile Organics
Apparent Losses
Total
Non Volatile Organics
Water and Non Volatile Organics
1648
1621
27
0
2830
3294
667
2527
1129
903
226
57.0
,104
12,
12.
0.
0
21.8
20.5
5.1
15.4
8.7
6.9
1.7
0.4
0.8
recorded. The effluent was a well mixed dispersion, which required overnight
standing to separate into two distinct phases. No insoluble oil phase
occurred. The input and yield data for this experiment are shown in TABLE
17.
Process 2B. Continuous Countercurrent Three Phase Extraction of Vacuum
Stripped Oil—
This experiment was conducted with'vacuum stripped oil by the same
method used in Process 2A. The inputs and yields for this experiment are
shown in TABLE 18.
The loss of nonvolatile organics is attributed to retention in the
solution remaining in the extractor at the end of the experiment. Evapora-
tion from the vigorously stirred system and minor leakage resulted in some
loss of solvents and volatiles.
65
-------
TABLE 17, INPUTS AND YIELDS. PROCESS 2A
Operation
Total g
g/minute
(average)
Inputs
Oil Sample In
Total Sample
Nonvolatile Organic
Volatile Organic
Sample Moisture
Extraction Solvent In
Water
MIBK
555
472
38
44
780
562
6.2
5.2
0.4
0.5
8.7
6.2
Outputs
Aqueous Phase
Nonvolatile Organics
Solvents and Volatile Organics
MIBK Phase
Nonvolatile Organics
Solvents and Volatile Organics
Insoluble Oil Phase
Apparent Losses
Nonvolatile Organics
Solvents and Volatile Organics
1153
358
795
741
114
627
None
0
3.0
12.8
4.0
8.8
8.2
1.3
7.0
0
~0
Extraction of Continuous Countercurrent Aqueous Phases
The water solution from Process 1A, Process IB, and Process 2B were
exhaustively extracted with successive small portions of MIBK. The aqueous
phase from Process 2A was not extracted with MIBK. The results of these
extractions and the results of the subsequent analyses of the phases and
fractions are summarized in TABLE 19.
j ..'.
The percent yields are expressed in terms of water-free oil including
volatile organics. The total nonvolatile organics recovered should approach
93 percent for unstripped oil and 97 percent for vacuum stripped oil. The
nonvolatile organics were approximately evenly distributed between the water
phase and the insoluble oil phase in Process 1A. The solubility of the
vacuum stripped oil was somewhat less than that of the unstripped oil.
In Process 2A the nonvolatile organics from the unstripped oil
appeared to concentrate in the aqueous phase, and the two phases separated
very slowly. These phases were stored for future applications research. In
66
-------
TABLE 18. INPUTS AND YIELDS. PROCESS 2B
g/tnlnute
Operation Total g (average)
Inputs
Oil Sample In
Total Sample 1,678 13.4
Nonvolatile Organics 1,629 13.0
Volatile Organics 49 0.4
Extraction Solvents In
Water 1,900 15.2
MIBK 1.198 9.6
Outputs
Aqueous Phase
Nonvolatile Organics
Solvents and Volatile Organics
MIBK Phase
Nonvolatile Organics
Solvents and Volatile Organics
Insoluble Oil Phase
Apparent Losses
Nonvolatile Organics
Solvents and Volatile Organics
2,746
735
'2,011
1,812
798
1,014
None
96
122
22.0
5.9
16.1
14.5
6.4
8.1
0.8
1.0
Process 2B , the nonvolatile organics were distributed almost equally between
the aqueous and MIBK phases. Whether in water extraction (Process 1) or
three phase extraction (Process 2) the presence of volatile organics enhanced
the water solubility of the nonvolatile organics.
The apparent yields and distributions of phenolic compounds also were
strongly dependent on the extraction method. In the water extraction experi-
ments (Process l) the total percent phenolic was apparently half of that
detected in the three phase extraction (Process 2) products. In the three
phase extraction of unstripped oil more than 80 percent of the phenolics were
concentrated in the aqueous phase. With stripped oil the phenolics were
distributed almost evenly between the aqueous and MIBK phases. The concen-
trations of polyhydroxy neutral compounds found in the water phases of
Process 1A, Process IB, and Process 2B are similar. The high apparent con-
centration of PNC in the Process 2 A water phase was believed to include some
neutrals of high aromaticity (NHA). The solubility of NHA compounds in water
was believed to be enhanced by the presence of volatile compounds from the
unstripped oil. This supposition was supported by the similarity of the sums
of PNC plus NHA in Process 2 A and Process 2B. Both sums are near 56 per-
cent.
67
-------
TABLE 19. COMPOSITION OF CONTINUOUS EXTRACTION YIELDS
Percent Percent Percent Percent
Process NVO* Phenolic PrtCf NHA**
Process 1A (Unstripped oil)
Aqueous Phase
Extracted Aqueous Fraction
MIBK Extract Fraction
Insoluble Oil Phase
Process IB (Vacuum stripped oil)
Aqueous Phase
Extracted Aqueous Fraction
MIBK Extract Fraction
Insoluble Oil Phase
Process 2 A (Unstripped oil)
Aqueous Phase
MIBK Phase
Process 2 B (Vacuum stripped oil)
Aqueous Phase
Extracted Aqueous Fraction
MIBK Extract Fraction
MIBK Phase
44.0
(33.4)
(10.6)
44.5
40.5
(33.1)
(7.4)
54.8
70.2
22.3
43.8
(35.7)
(8.1)
47.6
9.1 26.9
(6.5) (26.9)
(2.6) (-)
8.4
6.5 28.1
(5.0) (28.1)
(1.5)
7.9
31.4 38.7
5.9
16.1 24.4
(11.4) (24.4)
(4.7) (-)
17.6
8.0
(8.0)
36.1
5.9
(5.9)
46.9
16.4
3.4
(3.4)
30.0
* Non volatile organics
^ Polyhydroxy neutral compounds
**Neutrals of high aromaticity
68
-------
Vacuum Stripping of Oil
Moisture analyses of the oil samples by azeotropic distillation with
toluene indicated that about 14.7 percent of the sample was water and low
boiling water soluble compounds. Gas chromatography showed 8.2 percent
water and 6.5 percent volatile organics. These volatile materials could
represent a possible sample cut for separate processing and could also inter-
fere in the extraction of groups of higher molecular weight compounds in the
oil. Samples of the oil were vacuum stripped in a rotary evaporator at
three temperatures for varying lengths of time to determine the rate and
extent of volatiles removal. Results of these experiments are shown in
TABLE 20.
TABLE 20. VACUUM STRIPPING EXPERIMENTS
Temperature
(°c)
23.0
23.0
53.0
53.0
53.0
73.0
73.0
73.0
73.0
Time
(Hours )
40
60
0.5
1.0
4.0
0.4
0.7
1.0
2.5
P(min)
Torr
2
2
24
14
2
15
2
2
2
Percent Volatiles
Removed
13.7
13.8
8.2
11.2
16.4
9.8
13.8
15.0
18.7
The time required for vacuum stripping at 23°C was prohibitively long
for a continuous process. Heating the oil during vacuum stripping apparently
caused some chemical reactions, as the viscosity of the stripped oils
increased with both increasing time and temperature.
The percent of volatiles removed in these experiments is based on the
whole oil including volatiles but not water. The percent of volatiles
removed was calculated from the weight of the condensate in dry ice traps
between the evaporator and the vacuum pump. The thirty minute stripping
operation at 53°C was chosen as the basis for a semicontinuous stripping
operation to prepare oil samples for the continuous countercurrent extrac-
tions. The 8.2 percent volatiles removed included 5.1 percent water and 3.1
percent volatile organics by gas chromatography, and only minimal thickening
was observed in the stripped oil.
69
-------
Semicontinuous vacuum stripping experiments were carried out in
Buchler Model FE-2C* continuous rotary evaporators. The unstripped oil from
a calibrated reservoir wasqadmitted to the rotating evaporator bulb immersed
in a 53°C water bath and held under vacuum for 25 minutes before being
aspirated to a "stripped oil" reservoir. The distilled volatiles were col-
lected continuously in dry ice traps, and subsequently weighed and analyzed
by gas chromatography. The process was repeated using 200 ml portions of
unstripped oil until six liters of vacuum stripped oil had been collected.
The collected volatiles totalled 8.9 percent of the dry sample weight—5.9
percent water and 3.0 percent volatile organics.
Activated Carbon Adsorption Experiments
Three experiments were run contacting water extracts of pyrolytic oil
with activated carbon (Nuchar WV-G, Westvaco Carbon Co., Charleston, S.C.).
Slurry Contact with Stepwise Carbon Addition—
A 50 ml aqueous extract containing 15.9 g nonvolatile dissolved
organic material was stirred vigorously and treated with successive portions
of carbon until there was no further clarification of the color. After fil-
tering and washing the carbon with water, the combined filtrate and washings
were diluted with water to 100 ml. Evaporation of an aliquot portion of the
residing solution indicated that 8.9 g organics remained in solution and 7.0
g had been adsorbed on the carbon.
Elution of Aqueous Extract Through Activated Carbon—
A 10 ml portion of aqueous extract containing 2.9 g dissolved organics
was eluted through a 2.5 cm ID x 20 cm activated carbon column with water,
1:9 of methanol:water, 1:1 of methanol:water, methanol, and finally with
carbon disulfide. The eluted fractions were collected and evaporated to
dryness on a rotary vacuum evaporator. The results of this experiment are
summarized in TABLE 21.
TABLE 21. ORGANICS ELUTED FROM AQUEOUS CARBON COLUMN
Fraction
D-l
D-2
D-3
D-4
D-5
Eluting Solvent
Deionized Water
1:9 Methanol :Water
1:1 Methanol :Water
Methanol
Carbon Disulfide
ml
470
210
370
650
280
Organics
Eluted (g)
9
0.4405
0.2892
0.9968
0.5563
0.5557*
Total Organic
Eluted (g)
0.4405
0.7297
1.7265
2.2828
2.8385
* Eluted as small amount of very dark methanol phase and about 265 ml of
very pale carbon disulfide phase.
Buchler Instruments, Inc., Fort Lee, N. J.
70
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Inspection of TABLE 21 indicates that 15 percent of the organic
material eluted with water and nearly 63 percent eluted with methanol and
mixtures of methanol and water. The roughly 19 percent washed from the col-
umn with carbon disulfide was concentrated in a methanol layer on top of the
heavier carbon disulfide. Thus although carbon disulfide displaces the
adsorbed organic material left by methanol from the activated carbon, the
organic material is much more soluble in methanol than in carbon disulfide.
The fractions D-l through D-5 isolated in this experiment were analyzed by
TLC, LC and IR techniques, and the results were interpreted as follows. The
D-l fraction was shown to be quite polar from the TLG and LC reversed phase
column results. The IR spectra resembled the spectra of maltitol, an
alcohol carbohydrate. The results of the D-2 fraction were similar to those
of D-l. The data from TLC and LC with the D-3 fraction indicated the mater-
ial was polar, acidic and nonaromatic. The IR spectra resembled glyoxylic
acid. The TLC and LC results with D-4 indicated at least three polar compo-
nents were present, and the IR spectra of one of the components resembled
3-hydroxy-4-methoxyphenylethylene glycol.
Separation of Unstripped Oil on Activated Carbon—
A 50 g sample of unstripped oil was dissolved in methanol, and a 50
ml portion of the resulting solution containing 25.1 g of nonvolatile
organics was passed through a 2.5 cm ID x 50 cm carbon adsorption column,
previously prepared with a methanol-carbon slurry. The eluent in 30 ml por-
tions was returned to the top of the column until no further clarification
of the solution color was observed. The column was eluted with methanol (670
ml) until the eluted liquid was nearly colorless followed by elution with
210 ml carbon disulfide. The column was eluted then by 260 ml methanol which
was followed by a final elution with 100 ml of water. The results of these
elutions are summarized in TABLE 22, which show that 20 percent of the
organics were not eluted.
TABLE 22. ELUTION OF UNSTRIPPED OIL FROM ACTIVATED CARBON COLUMN
Elution Step
1-Methanol
2-Carbon Disulfide*
3-Methanol
4-Water
Solvent (ml)
560
210
260
100
Organic
Eluted (g)
13.12
4.99
1.60
0.40
Total Organic
Eluted (g)
13.12
18.11
19.71
20.11
* The eluent consisted of immiscible layers of methanol and carbon disulfide
and the organic material was concentrated in immiscible methanol layer.
Slurry Contact of Aqueous Extract of Pyrolytic Oil with Activated Carbon—
A 100 ml aliquot portion of a water extract from unstripped oil con-
taining 31.2 g dissolved nonvolatile organics was contacted with activated
carbon for 3 hours. Small quantities of the liquid were removed from the
71
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mixture at intervals, filtered and evaporated at 35°C on a vacuum evaporator.
These small samples indicated that the adsorption was nearly completed during
the first 10 minutes. After 3 hours the carbon was filtered from the solu-
tion, washed with water, and dried. The combined filtrate and washings were
evaporated to dryness in vacuo. The solutions contained 15.7 g organics
(50.3 percent of the organics in the sample). The dried carbon was exhaus-
tively extracted with N,N-dimethyl formamide (DMF), and the DMF extract con-
tained 11.1 organics (35.6 percent of the sample). The results show that
14.1% of the organics remained on the carbon.
Acid-Base Extraction of MIBK Phase with Ether
A portion of the MIBK solution from Extraction 11(4) which contained
phenolics, nonvolatile hydrocarbons, and neutrals of high aromaticity was
extracted with aqueous sodium hydroxide solution. The aqueous alkali
extract was extracted with diethyl ether and then acidified with dilute
sulfuric acid. The acidified solution was extracted with diethyl ether.
The phenolics in the final diethyl ether were determined by the NAT tech-
niques, and these results indicated that more than 90 percent of the phenolics
in the original MIBK phase had been extracted.
Fractional Distillation and Analysis of Fractions
• 'j
A 50 g sample of water-insoluble stripped oil, prepared by water
extraction of vacuum stripped oil, was vacuum distilled at approximately 6mm
in a short path simple column apparatus. The head temperatures and yields
are given in TABLE 23 below.
TABLE 23. DISTILLATION DATA FOR WATER-INSOLUBLE OIL
Fraction
No.
F-l
F-2
F-3
F-4
F-5
Residue
r,
Head Temperature
(°C)
50-100
100-110
110-120
120-175
175-193
Yields
(wt%)
10.5
6.3
4.6
7.9
17.3
53.4
The fractions F-l through F-5 were examined by several analytical
techniques to determine qualitatively the classes of the compounds and rela-
tive amounts. Thin layer chrotnatography indicated that F-l through F-4 con-
tained mainly two classes of compounds, phenolic aromatics and phenolic
ethers. TLC indicated that F-5 contained phenolic ethers, aromatic neutrals
and a trace of polyhydroxy neutral compounds. The analysis by liquid
chromatography confirmed TLC findings but yielded greater resolving power
among the phenolic compounds indicating F-l and F-2 had as many as 13
72
-------
compounds that were ultraviolet light absorbing. Infrared data indicate
predominantly phenolics and phenolic ethers in F-l through F-4 and aromatic
neutrals mixed with phenolic ethers in F-5.
Analytical Techniques
The identification of the different classes of organic functionality
has been accomplished by a variety of chemical analytical techniques. Our
immediate objective in this phase has been to rapidly determine the progress
of a separation process and identify the polyhydroxy compounds, dihydroxy
phenolics, phenolics, phenolic ethers and neutral aromatic classes in the
various phases or fractions. In most cases only a qualitative indication
was needed to complete the experiment since the phenolic components were
being determined by a NAT method. The fraction of neutrals of the sample was
determined by difference. To determine whether the neutral fraction was
primarily polyhydroxy aliphatic or aromatic or both, a TLC plate was run with
carbohydrate, phenolic and phenolic ether standards. To confirm these find-
ings an LC analysis at two wavelengths was made. An example of the results
obtained from all of the techniques applied to the three different phases
from a single process extraction is given in TABLE 24.
Nonaqueous Titration (NAT)—
A literature search was performed to determine suitable titration
methods for total phenolic material in the presence of carboxylic acids and
traces of water. Most of the conventional procedures utilized methods which
allowed only anhydrous conditions for determination. Based on the literature
search and experimental work, potassium hydroxide in methanol was chosen on
the basis of availability, ease of preparation and stability in storage. The
titration solvent chosen was dimethyl formamide because it has the required
basicity and compatibility with the water, neutral compounds and phenols
present in pyrolytic oils. DMF is relatively safe as compared to more vola-
tile amines and apparently yields adequate endpoint potentiometric millivolt
shifts.
Electrode systems were selected based on apparent end-point shifts in
millivolts on real pyrolytic oil samples. Both a glass calomel and platinum
versus platinum polarized electrode systems functioned adequately with known
standards which included acetic acid, benzoic acid, phenol and guaiacol.
However, the platinum polarized electrode system was the system that operated
best with real pyrolytic oil samples. Standardization was accomplished with
benzoic acid and guaiacol solutions, each 0.01N in DMF. The equipment used
to titrate samples was a semi-automatic recording titrimeter consisting of
the following components: (1) Pump - Cole Farmer Single channel, Variable
speed peristalic pump at 2.1 ml per minute; (2) Electrode - Platinum couple
Fisher Scientific K-F Titrimeter electrode; (3) Polarizer - Fisher Scientific
K-F Titrimeter Model 391; and (4) Recorder - Perkins Elmer Model 56.
The procedure for a determination was to standardize the semi-automatic
titration equipment with 3 ml samples of standard 0.01N benzoic acid and
0.01N guaiacol solutions in DMF. Oil samples for analysis were weighed in
the titrating vessel by difference. Each sample was titrated with the
methanolic potassium hydroxide solution until no further endpoints were noted.
73
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TABLE 24. ANALYTICAL RESULTS FROM BATCH EXPERIMENT PROCESS, l.A
Analytical
Technique
Aqueous Phase
extracted with MIBK
Phases
MIBK Extract of
Aqueous Phase
Insoluble Oil Phase
LC Predominately polar
polyhydroxy cpds; 3
dihydroxy phenolics
in moderate amts.
TLC Main components
polyhydroxy neutral
cpds with 3
dihydroxy phenolics
NAT 6.5% phenolic
27% polyhydroxy
neutrals
Predominately
phenols, dihydroxy
phenolics; trace
of polyhydroxy cpds
Three phenolic
cpds; only trace
amts. of poly-
hydroxy neutrals
2.9% phenolic
6.2% neutrals
Predominately
aromatic neutrals;
moderate amt. of
phenolics and trace
of polyhydroxy cpds
Strongly aromatic
neutral components;
moderate phenolic
content; no trace
of polyhydroxy cpds
8.4% phenolic
36% neutrals
Strong hydroxy func-
tionality; strong
ether functionality;
weak phenolic (,
indications
Indicated strong
phenolics and
ethers
Aromatic ketones;
subt'd aromatics;
phenolic
GC Only small amount of
sample eluted,
approx. 80% of sample
coked in the
inj ector
GC Silylation of sample
produced three irregu-
lar peaks of high
boiling character
similar to sugar cpds
Many phenolic and
creslyic cpds
Some phenols; ether
phenols
In calculating the results, it was assumed that the average molecular weight
of the phenolics was 125 and of the carboxylic acids, 100.
Thin Layer Chromatography—
Thin layer chromatography (TLC) offered an analytical technique which
could supplement the other techniques used in this study, particularly HPLC.
A separation by TLC of the general classes include the polyhydroxy
74
-------
carbohydrates, dihydroxy phenolics, phenolics, ether phenolics and aromatic
neutrals. The TLC separations were carried out with EM Silica Gel 60F-254
plates, 20 x 20 cm, and the solvent systems and detection (visualization)
reagents are given in TABLE 25.
TABLE 25. TLC SOLVENTS AND DETECTION REAGENTS
S-l"
Solvent Systems
S-2
S-3
Ethyl acetate
Acetonitrile
Water
65
25
10
N-butanol
Acetone
Water
40
50
10
Methanol
Benzene
Water
14
79
7
D-l
D-2
Detection Reagents
D-3
D-4
Bial's Orcinol
reagents used at
110 °C for 5 min.
Sulfuric acid
and potassium
dichromate
charring at 160°
for 10 min.
Ultraviolet
light at
254 and
365 nm
Diazotized
R Salt.
Scarlet
* The numbers after each solvent represents the percent by volume of each
solvent in the three component system.
The general procedure for a TLC analysis was as follows. The TLC
plates were normally activated for 15 min in a HO°C oven. Three microliter
samples 10 mg/ml in acetonitrile were applied. Each spot was dried and the
plate was developed in a presaturated tank of a chosen solvent system. After
a 10 to 14 cm rise of solvent the plate was dried in a low heat oven 80°C for
10 minutes and visualized with the detection agent of choice. Inspection by
U V light was usually done before any chemical reagent was applied. Rf values
were calculated by conventional means using the solvent front as R£ 100 and
the spotting point as Rj 0. Interpretation of the chromatograms was made
using standard compounds when possible and color reactions of the various
visualization reagents.
Gas Chromatography—
Gas chromatography as an analytical method was used almost exclusively
in Phase III of this project to analyze the volatiles fraction, obtained from
the vacuum stripping separation process. The conditions used for these
analyses were:
75
-------
Column 1. Pora Pak Q, 270 cm x 0.31 cm S.S.; oven, 180°C; injector, 200°C;
thermal conductivity detection, 175 ma; Helium carrier at 20
ml/min
Used for the determination of water, methanol, formic acid,
acetic acid, and propionic acid.
Column 2. SP-2100, 10% on HMDCS treated 100-120 mesh Supelcoport;
300 cm x 0.31 cm S.S.; FID; N_ carrier at 20 ml/min; oven
60°C; injector 100°C;
Used for the determination of furfural.
Infrared Spectroscopy—
Infrared spectra were made of the various fractions obtained in
the experiments with the continuous extraction processes. The spectra were
found to contain only fragmentary information due to the multiplicity of
compounds in each fraction. The overlapping of peaks precluded interpreta-
tion in only but the most general terms. Main bands of interest used in
this program are given in TABLE 26.
TABLE 26. INFRARED BANDS
Micron Wavelength Description
3.0 Broad hydrogen bonded OH function
3.8 Shoulder of carboxylic acid OH stretching
5.85 - 5.95 Carbonyl absorption
6.25 - 7.35 Carboxylate anion absorptions
10, 11, 7.1 Vinyl group absorptions
6.24, 12-14 Aromatic absorption bands
Liquid Chromatography—
Conditions used for LC analysis of fractions of the oils in this
phase of the project are given in TABLE 27.
TABLE 27. LIQUID CHROMATOGRAPHY CONDITIONS
Item Condition
Column: Spherosol ODS CTO 25 cm
Solvent Gradient: 0 - 100% linear. Total time 60 min,
S.olvent: 0 - 100% Acetonitrile in water
Detection: 254 nm, 190 nm
Sensitivity: 0.2 absolute
Chart speed: 8 inches per hour
76
-------
SECTION 6
PILOT PLANT DESIGN
PROCESS DESCRIPTION
Pyrolysis oils contain four classes of organic compounds in addition
to water which is condensed along with the organics. The classes are:
phenolics, neutrals of high aromaticity (NHA), polyhydroxy aromatics, acids,
and water. Separation work at the bench level led to the development of four
individual processing schemes:
Process 1-A 2 Stage continuous extraction—raw oil
Process 1-B 2 Stage continuous extraction—vacuum stripped oil
Process 2-A Continuous, simultaneous extraction—raw oil
Process 2-B Continuous, simultaneous extraction—vacuum stripped oil
Flow sheets for the four separation process and for the combined pilot plant
system are presented in Figures 39 through A3. A component-by-component
description of the processes follows.
Process 1-A 2 Stage Continuous Extraction—Raw Oil
Starting at the left in Figure 39 , raw pyrolytic oil, received in
barrels, is pumped into the raw oil storage tank (1). In preparation for a
processing run the raw oil is pumped into the raw oil feed tank (2). The raw
oil feed tank is equipped with a stirrer or mixer, to provide a homogeneous
feedstock. As the ambient temperature decreases the pyrolytic oil becomes
more viscous. A recycle loop with a heating device is included to raise the
temperature of the pyrolytic oil into the 100-130°F range, as necessary, to
provide the proper flow of oil.
Raw pyrolytic Oil is pumped into the extractor (3) above the mixers
(near the top of the extractor). Water from the water storage tank (4) is
pumped into the extractor at approximately the same height as the pyrolytic
oil (and above the mixers). Two or more mixers provide violent agitation and
intimate mixing of the pyrolytic oil and water. A recycle stream draws a
portion of the oil-water mixture from approximately the height of the mixers
and returns the mixture to the extractor near the bottom, above the level of
the spent, insoluble oil. The recycle line is equipped with a heating
device to raise the temperature of the mixture from ambient to about 150°F,
as is necessary. Spent oil droplets descend through the extractor and
accumulate in the bottom of the extractor. Excess spent, insoluble oil is
pumped to the spent oil storage tank (6), while always maintaining a level
of spent oil in the extractor.
77
-------
oo
Fan Oil
Storage Tank
Preheater
n
Raw Oil
Feed Tank
(7)
Water Soluble
/
lup
Extractor
Water
Storage Tank
ank (
Ho]
\
{
(9)
„
>-
MTBK Soluble
~1
Ldup Tank JF
f^\
(10)
~-~
Extractor
^\
(8)
tfl
Water
Soluble
^ \ .
-v, U3) y ^
To
Recycle
Holdup Tank
(14)
M1BK
MIBK Soluble
Column or
Evaporator
M1BK
Storage Tank
Spent Oil
Storage Tank
Water
lnsolubl<
Water
Soluble
(15)
N x Product
Vacuum Storage Tank
Evaporator
•Storage Tank
Figure 39. Separation process no. lA--raw oil-—2 stage extraction
-------
The stream of water and soluble orgaiiics exits near the top of the
extractor and is pumped into a holdup tank (7). The material in the holdup
tank is pumped into the 2nd stage extractor (8) at a level above the mixers.
Methyl isobutyl ketone (MIBK) is pumped into the extractor from the MIBK stor-
age tank (5), at approximately the same level as the water soluble organic
inlet stream. Two or more mixers or stirrers provide violent agitation and
intimate mixing of the water soluble organics and the MIBK. A recycle stream
draws a portion of the water soluble organic—solvent mixture from the level
of the mixers and returns the mixture to the extractor near the bottom. A
phase separation occurs in the extractor, with the heavier aqueous solution
settling to the bottom of the extractor, and the lighter organic solution
moving toward the top of the extractor. The aqueous solution is removed from
the extractor near the bottom and is pumped into the water soluble holdup
tank (13), and then into the vacuum evaporator (14) where the water soluble
organics are separated from the water. The water is vaporized and returned
to,the water storage tank (4). The organics are pumped into the water soluble
organics—product storage tank (15).
The organic phase from the second stage extractor exits near the top
of the extractor and is pumped into the MIBK soluble holdup tank (9). The
organic phase is then fed into an evaporator (or column) (1Q) where the MIBK
is vaporized and collected in the MIBK-holdup tank (11). The recovered MIBK
is then returned to the MIBK storage tank (5). The MIBK soluble organics are
concentrated in the evaporator and pumped to the MIBK soluble organics—pro-
duct storage tank (12).
Process 1-B 2 Stage Continuous Extractor—Vacuum Stripped Oil
Starting at the left in Figure 40, raw pyrolytic oil, received in
barrels, is pumped into the raw oil storage tank (1). In preparation for a
processing run the raw oil is pumped into the raw oil feed tank (2) . The
raw oil feed tank is equipped with a stirrer or mixer, to provide a homogene-
ous feed stock. A recycle loop with a heating device is included to raise
the temperature of the pyrolytic oil into the 100-130°F range, as necessary,
to provide the proper flow of oil.
The raw pyrolytic oil is pumped into a vacuum evaporator (or vacuum
stripping column) (3), to remove the volatiles. The volatiles are components
that are vaporized at atmospheric pressure at 100°F (212°F). They consist
of water (60-70%), acetic acid ( ~ 20%) and small amounts of other low boiling
organic compounds. The volatiles are condensed and pumped to the volatiles
storage tank (4). The vacuum stripped oil is pumped from the stripper to the
1st stage extractor (5), and enters the extractor above the mixers (near the
top of the extractor). Water from the water storage tank (6) is pumped into
the extractor at approximately the same height as the vacuum stripped oil
(and above the mixers). Two or more mixers provide violent agitation and
intimate mixing of the vacuum stripped oil and water. A recycle stream draws
a portion of the oil-water mixture from approximately the height of the
mixers and returns the mixture to the extractor near the bottom, above the
level of the spent insoluble oil. The recycle line is equipped with a heating
device to raise the temperature of the mixture from ambient to about 150°F,
79
-------
oo
O
Raw Oil
Storage Tank
Preheater
n
Condenser
Volatiles
(9)
Raw Oil
Feed Tank
Vacuum Stripper or
Vacuum Evaporator
(3)
!3_1
Storage
Tank
Reheateri
Water Soluble
I
1
1
1
c
ly
->-
A
n
Q-
(5)
— L
r—
— L
1
1.
c*>
T
« * 1
Recycle
Extractor
To
Recycle
f MIBK Soluble
Tank C
Hoi
i
(11)
>-
~l
dup Tank Jf
/•-ION
Extractor
fK
(10)
^
Water
. . . .-^
Soluble
^/ /I C\ \ _
(16)
Holdup Tank
MIBK
Holdup Tank
Column or
Evaporator
~»i MIBK Soluble
-W
Water
Storage Tank
MIBK
Storage Tank
Spent Oil
Storage Tank
Water
Insoluble
(14)
Water
Soluble
(17)
/
Product
Vacuum Storage Tank
Evaporator
Storage Tank
Figure 40. Separation process IB—vacuum stripped—2 stage extraction
-------
as is necessary. Spent oil droplets descend through the extractor and accum-
ulate in the bottom of the extractor. Excess, spent insoluble oil is pumped
to the spent oil storage tank (8), while always maintaining a level of spent
oil in the extractor.
The stream of water and soluble organics exits near the top of the
extractor and is pumped into a holdup tank (9). The material in the holdup
tank is pumped into the 2nd stage extractor (10) at a level above the mixers.
Methyl isobutyl ketone (MIBK) is pumped into the extractor from the MIBK stor-
age tank (7), at approximately the same level as the water soluble organic
inlet stream. Two or more mixers or stirrers provide violent agitation and
intimate mixing of the water soluble organics and the MIBK. A recycle stream
draws a portion of the water soluble organic—solvent mixture from the level
of the mixers and returns the mixture to the extractor near the bottom. A
phase separation occurs in the extractor, with the heavier aqueous solution
settling to the bottom of the extractor, and the lighter organic solution
moving toward the top of the extractor. The aqueous solution is removed from
the extractor near the bottom and is pumped into the water soluble holdup tank
(15), and then into the vacuum evaporator (16) where the water soluble organics
are separated from the water. The water is vaporized and returned to the
water storage tank (6). The organics are pumped into the water soluble organ-
ics—product storage tank (17).
The organic phase from the second stage extractor exits near the top
of the extractor and is pumped into the MIBK soluble-holdup tank (11). The
organic phase is then fed into an evaporator (or column) (12) where the MIBK
is vaporized and then collected in the MIBK-Holdup tank (13). The recovered
MIBK is then returned to the MIBK storage tank (7) . The MIBK soluble organics
are concentrated in the evaporator and pumped to the MIBK soluble organics—
product storage tank (14).
Process 2-A Continuous, Simultaneous Extraction—Raw Oil
Starting at the left in Figure 41, raw pyrolytic oil, received in
barrels,-is pumped into the raw oil storage tank (1). In preparation for a
processing run the raw oil is pumped into the raw oil feed tank (2). The raw
oil feed tank is equipped with a stirrer or mixer, to provide a homogeneous
feedstock. As the ambient temperature decreases the pyrolytic oil becomes
more viscous. A recycle loop with a heating device is included to raise the
temperature of the pyrolytic oil into the 100-130°F range, as necessary, to
provide the proper flow of oil.
Raw pyrolytic oil is pumped into the extractor (3) above the mixers
(near the top of the extractor). Water from the water storage tank (4) is
pumped into the extractor at approximately the same height as the pyrolytic
oil (and above the mixers). MIBK is pumped into the extractor from the MIBK
storage tank (5) at a level below the mixers. Two or more mixers provide
violent agitation and intimate mixing of the pyrolytic oil, water, and MIBK.
A recycle stream draws a portion of the oil-water-MIBK mixture from
approximately the height of the mixers and returns the mixture to the extractor
near the bottom. The recycle line is equipped with a heating device to raise
the temperature of the mixture from ambient to about 150°F, as is necessary.
81
-------
oo
Raw Oil
Storage Tank
Reheater
-. r.
Raw Oil
Feed Tank
(6)
MIBK Soluble
Separator
(7)
Holdup Tank
ft 2.
©• ( <»> )
To
Recycle
MIBK
Holdup Tank
(8)
Extractor
Column or
Evaporator
(_UO) )
Q+°
(3)
Recycle
Extractor
Water
Soluble
(11)
Holdup
Tank
(12)
Vacuum
Evaporator
Storage Tank
Figure 41. Separation process 2A—raw oil—simultaneous extraction.
Water
Storage Tank
(5)
MIBK
Storage Tank
Water
Soluble
(13)
Product
Storage Tank
-------
The oll-water-MIBK mixture exits near the top of the extractor and is
pumped into a separator (6). A phase separation occurs, with the heavier
aqueous solution settling to the bottom, and the lighter organic solution
moving toward the top of the separator. The aqueous solution is pumped from
the separator to the water soluble holdup tank (11), and then into the vacuum
evaporator (12) where the water soluble organics are separated from the water.
The water is vaporized and returned to the water storage tank (4). The
organics are pumped into the water soluble organics—product storage tank (13).
The organic phase from the separator exits from the top of the separa-
tor and is pumped into the MIBK soluble-holdup tank (7). The organic phase
is then fed into an evaporator (or column) (8) where the MIBK is vaporized
and collected in the MIBK holdup tank (9). The recovered MIBK is then
returned to the MIBK storage tank (5). The MIBK soluble organics are concen-
trated in the evaporator and pumped to the MIBK soluble organics—product
storage tank (10).
Process 2-B Continuous, Simultaneous Extraction—Vacuum Stripped Oil
Starting at the left in Figure 42 , raw pyrolytic oil, received in
barrels, is pumped into the raw oil storage tank (1). In preparation for a
processing run the raw oil is pumped into the raw oil feed tank (2). The raw
oil feed tank is equipped with a stirrer or mixer, to provide a homogeneous
feedstock. A recycle loop with a heating device is included to raise the
temperature of the pyrolytic oil into the 100-130°F range, as necessary, to
provide the proper flow of oil.
The raw pyrolytic oil is pumped into a vacuum evaporator (or vacuum
stripping column) (3), to remove the volatiles. The volatiles are components
that are vaporized at atmospheric pressure at 100°F (212°F). They consist of
water (60-70%), acetic acid ( S20%) and small amounts of other low boiling
organic compounds. The volatiles are condensed and pumped to the volatiles
storage tank (4). The vacuum stripped oil is pumped from the stripper to the
extractor (5), above the mixers (near the top of the extractor). Water from
the water storage tank (6) is pumped into the extractor at approximately the
same height as the vacuum stripped oil (and above the mixers). MIBK is pumped
into the extractor from the MIBK storage tank (7) at a level below the mixers.
Two or more mixers provide violent agitation and intimate mixing of the vacuum
stripped oil, water and MIBK. A recycle stream draws a portion of the oil-
water-MIBK from approximately the height of the mixers and returns the mixture
to the extractor near the bottom. The recycle line is equipped with a heating
device to raise the temperature of the mixture from ambient to about 150°F,
as is necessary.
The vacuum stripped oil-water-MIBK mixture exits near the top of the
extractor and is pumped into a separator (8). A phase separation occurs, with
the heavier aqueous solution settling to the bottom, and the lighter organic
solution moving toward the top of the separator. The aqueous solution is
pumped from the separator to the water soluble holdup tank (13), and then into
the vacuum evaporator (14) where the water soluble organics are separated from
the water. The water is vaporized and returned to the water storage tank (6).
83
-------
00
OED
Condenser
Volatiles
Raw Oil
Storage Tank
Preheater
«-
(8)
±
Raw Oil
Feed Tank ^r
Vacuum Stripper or
Vacuum Evaporator
Recycle , Extractor
To MIBK
Recycle Holdup
Column or
Evaporator
Water
Storage Tank
MIBK
Storage Tank
Storage Tank Evaporator
Figure 42. Separation process 2B—vacuum stripped—simultaneous extraction.
Water
Soluble
, Product
Storage Tank
-------
The organics are pumped into the water soluble organics - product storage tank
(15).
The organic phase from the separator exits from the top of the separator
and is pumped into the MIBK soluble-holdup tank (9) . The organic phase is
then fed into an evaporator (or column) (10), where the MIBK is vaporized and
collected in the MIBK holdup tank (11). The recovered MIBK is then returned
to the MIBK storage tank (7). The MIBK soluble organics are concentrated in
the evaporator and pumped to the MIBK soluble organics-product storage tank
(12).
DESIGN OF THE PILOT PLANT
The pilot plant processing scheme, in each of the four cases is based
on the results of the batch and continuous extraction data produced during
the laboratory experiments. Pilot equipment must be large enough to provide
the data necessary to accurately scale up to design a commercial pyrolytic oil
processing plant. But a major constraint on the size of the pilot plant is
the availability of the pyrolytic oil. The basis for the sizing of the pilot
plant is given below. (See Appendices A and B for calculations.)
Size of Pilot Plant
The proportions of the input and output streams for the four cases of
the pilot plane design are determined by the rate data provided by the lab
scale continuous extraction data. To arrive at a flowrate for pilot plant
use, the actual residence time for each extraction method is inspected. TABLE
28 shows the input rates in grams per minute and ml per minute for each
component, the total volume flowrate in ml per minute, the extractor volume
in ml , and the residence time in minutes.
TABLE 28
INPUT RATES TO EXTRACTOR
Total
Input Rate
grains /min
Process Oil Water
1-A
1-B
2-A
2-B
10.
13.
6.
13.
60
10
19
28
22.34
23.30
9.39
15.54
Input Rate
ml /min
MIBK Oil Water
8
— 10
6.18 5
9.75 10
.72
.58
.02
.73
4
22.34
23.30
9.39
15.54
Input
Rate
MIBK ml /min
31
33
7.72 22
12.17 38
.06
.88
.13
.44
Extractor
Volume
ml
1950
1950
1950
1950
Residence
Time
Min.
62.
57.
88.
50.
77
56
13
73
Choosing a minimum residence time of 65 minutes, TABLE 29 shows the required
extractor volume for various oil input rates and total input rates.
85
-------
TABLE 29. REQUIRED EXTRACTOR VOLUME
Input
Rate
GPM
3
4
5
6
Total Input
Rate
GPM
9
12
15
18
Extractor
Volume
Gal.
588
780
975
1170
Minimum
Oil Input Rate
Gal.
195
260
325
390
As calculated from TABLE 29 the oil input rate (in ml per minute) is approxi-
mately 1/3 of the total input rate (ml per min). The oil input rate for the
pilot plant design was selected to be 4 GPM.
COST SUMMARY
The costs for the major equipment necessary to conduct tests using any
of the 4 processing schemes is shown in TABLE 30. The total installed equip-
ment cost is $365,900. The pilot plant equipment cost including instrumenta-
tion and controls, electrical, and piping is $508,000. (See Appendix B).
TABLE 30. PILOT PLANT - COST SUMMARY
Raw Oil Storage Tank
Raw Oil Feed Tank
Vacuum Evaporator (Stripper)
Extractor (1st Stage)
Separator (or Holdup Tank
Extractor (2nd Stage)
MIBK Soluble - Holdup Tank
Evaporator
MIBK Holdup Tank
MIBK Soluble - Product Storage Tank
Water Soluble - Holdup Tank
Vacuum Evaporator
Water Soluble - Product Storage Tank
MIBK Storage Tank
Volatiles - Product Storage Tank
Spent Oil - Product Storage Tank
Water Storage Tank
Total Installed Equipment Cost
(1)
(2)
(3)
(4)
(5)
(6)
(7)
(8)
(9)
(10)
(11)
(12)
(13)
(14)
(15)
(16)
(17)
$9,382
9,382
39,090
48,790
23,454
48,790
9,382
46,908
3,440
4,691
9,382
87,561
7,193
3,440
4,691
4,691
5,629
365,900
Instrumentation and Controls (9.35% of installed
equipment cost) 34,200
Piping - (22.3% of installed equipment cost) 81,600
Electrical - (7.2% of installed equipment cost) 26,300
Total Pilot Plant Equipment Cost $508,000
86
-------
00
d IK
Raw Oil
Storage Tank
Preheater
•«-
±
Raw Oil
Feed Tank
Vacuum Stripper
or
Vacuum Evaporator
S
(L_D
d
Condenser
Water
Storage Tank
,. /
Separator or Holdup
Holdup Tank Tank
o
To MIBK
Recycle Holdup Tank
Column or
Evaporator
Storage Tank
Vacuum
Evaporator
MIBK
Storage Tank
Spent Oil
Storage Tank
Water
Soluble
Product
Storage Tank
Figure 43. Pyrolysis oil pilot plant schematic—continuous process.
-------
SECTION 7
DESIGN AND ECONOMICS OF COMMERCIAL SIZE PLANT
PROCESS DESCRIPTION
The four processing schemes examined at the pilot plant scale are fur-
ther investigated at the size of a commercial facility. A component by com-
ponent description of the four processes is given in Section 6. Flow sheets
for the four separation processes are shown in Figures 39 through 42. At this
preliminary stage of the process design, the commercial size plant and the
pilot plant differ only in the size of the major process equipment. The pro-
cess descriptions remain the same.
DESIGN BASIS
The full scale, commercial pyrolytic oil processing plant is based on
the availability of pyrolytic oil. The oil will be provided by one or more
wood pyrolysis plants. It is possible that future pyrolytic oil processing
plants will use oil produced from sources other than wood - other agricultural
or cellulosic materials, or municipal refuse - but for the purposes of this
design only wood pyrolysis will be considered.
Georgia Tech has had considerable experience with the Georgia Tech -
Tech-Air Corporation pyrolysis system, which produces char and pyrolytic oil
by the pyrolysis of wood. Although other processes are available to produce
pyrolytic oil, no other process has performed reliably on a continuous basis
over an extended period of time. The Tech-Air Corporation has operated a
pyrolysis plant, using the Georgia Tech - Tech-Air process over a period of
several years in South Georgia. That plant had a nominal processing rate of
1-1/2 to 2 tons per hour of dried wood material. In addition the Stanford
Research Institute (SRI) stated that the Georgia Tech - Tech-Air technology
was the closest to commercialization of all the processes investigated [24].
Therefore, the Georgia Tech Pyrolysis Process [6] will be used as a basis for
the supply of pyrolytic oil.
Preliminary design calculations have been made to scale the Georgia
Tech pyrolysis process up to anywhere from 3.5 tons per hour to several hun-
dred tons per hour, based on a dried wood feed material. SRI uses a plant
size of 1,000 ton per day , dry wood feed rate, or approximately 42 tons per
hour. The SRI study used four 10 ton per hour (dry feed rate) pyrolyzers
operating in parallel.
Since the size of the largest operating pyrolysis plant to date is only
1-1/2 to 2 tons of dried feed per hour, it is not likely that the next
88
-------
generation of pyrolyzers will be scaled up to 10 tons per hour. An interme-
diate size in the range of 3.5 to 7 tons per hour will probably be built and
tested for a period of time. It is estimated that the data currently avail-
able will permit the construction of a nominal 5 ton per hour pyrolyzer with
limited risk regarding performance. It has been projected that 5 ton per hour
pyrolyzer is large enough to adequately provide a return on the capital
investment, while minimizing the unknowns associated with scale up.
Therefore it is projected that five 5 ton per hour (dry feed rate)
pyrolysis plants will provide oil to the pyrolytic oil processing plant. The
pyrolysis plants are estimated to operate with an 18% oil yield based on the
dry feed rate to the pyrolyzers, with an operating year of 345 days. Thus the
oil processing plant must be located in proximity to 25 tons per hour of
pyrolysis processing capacity. This requirement is conservative when compared
to the SRI scenario of individual pyrolysis plants of 42 tons per hour. Based
on the conditions above, one 5 ton per hour pyrolysis plant will produce
14,904,000 pounds of oil per operating year or 1,419,400 gallons per year.
The combined output of the 5 pyrolysis plants is 74,520,000 pounds per year or
7,097,000 gallons per year.
ECONOMICS AND FEASIBILITY
The economic feasibility of each process discussed in Section 6 has
been considered for a commercial size plant. This analysis included total
capital investment with equipment costs, manufacturing and product costs,
depreciation and estimated income.
Itemized equipment posts and equipment sizing calculations are included
in Appendix C. Each of the processes (1A, IB, 2A, 2B) were treated as a sep-
arate case. Cost summaries for the major equipment for each of the processes
are~given in Appendix C and in TABLES 31-34. The total installed equipment
costs for each of the processes are: 1A, $1,127,000; IB, $1,172,000; 2A,
$1,025,000; and 2B, $1,036,000. The equipment costs are included in the
direct costs of the total capital investment.
The total capital investment, which included direct and indirect costs
and working capital, were calculated for each process and are summarized on
pages^93-96. The manufacturing and total product costs which include raw
materials, labor, utilities, maintenance, operating supplies, laboratory costs
and direct production costs are summarized on pages 97-100. Depreciation is
discussed on pages 101-102.
In order to arrive at estimated current selling prices for potential
chemical fractions from pyrolytic oil, prices for similar organic substances
were selected and used from the Chemical Marketing Reporter of December, 1979.
Income,-,was calculated for a minimum, average and maximum selling price, and
thes^ results are summarized on pages 102-107. The rate of return analysis
is presented on pages 107-116.
89
-------
TABLE 31. PROCESS 1A - 2 STAGE CONTINUOUS EXTRACTION -
RAW OIL - INSTALLED EQUIPMENT COST SUMMARY
Raw Oil Storage Tank
Raw Oil Feed Tank
Extractor - 1st Stage
Water Storage Tank
MIBK Storage Tank
Spent Oil Storage Tank
Holdup Tank
Extractor - 2nd Stage
MIBK Soluble - Holdup Tank
Evaporator
MIBK Holdup Tank
MIBK Soluble - Product Storage Tank
Water Soluble - Holdup Tank
Vacuum Evaporator
Water Soluble - Product Storage Tank
Total Installed Equipment Cost
(1)
(2)
(3)
(4)
(5)
(6)
(7)
(8)
(9)
(10)
(ID
(12)
(13)
(14)
(15)
$187,600
9,400
91,500
31,300
8,100
136,100
39,100
99,100'
31,300
71,300
31,J300
50,000
41,900
187,600
111,000
$1,126,600
TABLE 32. PROCESS IB - 2 STAGE CONTINUOUS EXTRACTION -
VACUUM STRIPPED OIL-INSTALLED EQUIPMENT COST SUMMARY
Raw- Oil Storage Tank
Raw 'Oil Feed Tank
Vacuum Evaporator - Raw Oil
Volatiles Storage Tank
Extractor - 1st Stage
Water Storage Tank
MIBK Storage Tank
Spent Oil Storage Tank
Holdup Tank
Extractor - 2nd Stage
MIBK Soluble - Holdup Tank
Evaporator
MIBK Holdup Tank
MIBK Soluble - Product Storage Tank
Water Soluble - Holdup Tank
Vacuum Evaporator
Water Soluble - Product Storage Tank
Total Installed Equipment Cost
(1)
(2)
(3)
(4)
(5)
(6)
(7)
(8)
(9)
(10)
(11)
(12)
(13)
(14)
(15)
(16)
(17)
$187,600
9;400
78,200
73,500
77,800
46,900
7,500
139,500
36,000
80,800
29,100
51,600
28,100
39 , 100
36; 900
150,100
100,000
$1,172,100
90
-------
TABLE 33. PROCESS 2A - CONTINUOUS, SIMULTANEOUS EXTRACTION -
RAW OIL - INSTALLED EQUIPMENT COST SUMMARY
Raw Oil Storage Tank
Raw Oil, Feed Tank
Extractor
j >
Water,,; Storage Tank
MIBK Storage Tank
Separator
MIBK Soluble - Holdup Tank
Evaporator
MIBK Holdup Tank
MIBK-.Soluble - Product Storage Tank
Water. Soluble - Holdup Tank
Vacuum* Evaporator
Water Soluble - Product Storage Tank
,. Total Installed Equipment Cost
(1)
(2)
(3)
(4)
(5)
(6)
(7)
(8)
(9)
(10)
(11)
(12)
(13)
$187,600
9,400
120,400
46,900
8,100
55,400
36,600
100,100
31,300
79,400
40,700
156,400
153,200
$1,025,500
TABLE 34. PROCESS 2B - CONTINUOUS, SIMULTANEOUS EXTRACTION -
VACUUM STRIPPED OIL - INSTALLED EQUIPMENT COST SUMMARY
Raw Oil Storage Tank
Raw Oil Feed Tank
Vacuum Evaporator
Volat,iles Storage Tank
Extractor
Water, Storage Tank
MIBK "Storage Tank
Separator
MIBK Soluble - Holdup Tank
Evaporator
MIBK Holdup Tank
MIBK "Soluble - Product Storage Tank
Water Soluble - HOldup Tank
Vacuum Evaporator
Water Soluble - Product Storage Tank
Total Installed Equipment Cost
(1)
(2)
(3)
(4)
(5)
(6)
(7)
(8)
(9)
(10)
(ID
(12)
(13)
(14)
(15)
$187,600
9,400
78,200
73,500
83,900
37,500
6,900
41,900
28,100
65,700
23,400
111,000
36,000
134,500
118,800
$1,036,400
91
-------
Total Capital Investment
Direct Costs—Process 1-A 2 Stage Continuous Extraction—Raw Oil
Purchased Equipment—Installed (End '79) $1,126,600
Instrumentation and Controls—Installed 105,300
- 9.35% of Installed Equipment Costs
Piping—Installed 251,200
- 22.3% of Installed Equipment Costs
Electrical—Installed 81,100
- 7.2% of Installed Equipment Costs
Buildings—Including Services 234,300
- 20.8% of Installed Equipment Costs
Yard Improvements 81,100
- 7.2% of Installed Equipment Costs
Service Facilities—Installed 446,100
- 39.6% of Installed Equipment Costs
Total Direct Plant Cost $2,325,700
Indirect Costs—
Engineering and Supervision $ 283,900
- 25.2% of Installed Equipment Costs
Construction Expense 235,500
- 20.9% of Installed Equipment Costs
Total Direct and Indirect Costs $2,845,100
Contractor's Fee 142,300
- 5% of Direct and Indirect Costs
Contingency 284,500
- 10% of Direct and Indirect Costs
Fixed Capital Investment 3,271,900
Working Capital 363,500
- 10% of Total Capital Investment
Total Capital Investment $3,635,400
92
-------
Direct Costs—Process IB—2 Stage Continuous Extraction—Vacuum Stripped Oil
Purchased Equipment—Installed (End '79) $1,172,100
Instrumentation and Controls—Installed 109,600
- 9.35% of Installed Equipment Costs
Piping—Installed 261,400
- 22.3% of Installed Equipment Costs
Electrical—Installed 84,400
- 7.2% of Installed Equipment Costs
Buildings—Including Services ; 243,800
- 20.8% of Installed Equipment Costs
Yard Improvements 84,400
- 7.2% of Installed Equipment Costs
Service Facilities—Installed 464,100
- 39.6% of Installed Equipment Costs
Total Direct Plant Cost— 2,419,800
Indirect Costs—
Engineering and Supervision 295,400
- 25.2% of Installed Equipment Costs
Construction Expense 244,900
- 20.9% of Installed Equipment Costs
Total Direct and Indirect Costs 2,960,100
Contractor's Fee 148,000
- 5% of Direct and Indirect Costs
Contingency 296,000
- 10% of Direct and Indirect Costs
Fixed Capital Investment 3,404,100
Working Capital 378,200
- 10% of Total Capital Investment
Total Capital Investment ' $3,782,300
93
-------
Direct Costs—Process 2A—Continuous, Simultaneous Extraction—Raw Oil
Purchased Equipment—Installed (End '79) $1,025,500
Instrumentation and Controls—Installed 95,900
- 9.35% of Installed Equipment Costs
Piping—Installed 228,700
- 22.3% of Installed Equipment Costs
Electrical—Installed 73,800
- 7.2% of Installed Equipment Costs
Buildings—Including Services 213,300
- 20.8% of Installed Equipment Costs
Yard Improvements 73,800
- 7.2% of Installed Equipment Costs
Services Facilities—Installed 406,100
- 39.6% of Installed Equipment Costs
Total Direct Plant Cost— $2,117,100
Indirect Costs
Engineering and Supervision 258,400
- 25.2% of Installed Equipment Costs
Construction Expense 214,300
- 20.9% of Installed Equipment Costs
Total Direct and Indirect Costs 2,589,800
Contractor's Fee 129,500
- 5% of Direct and Indirect Costs
Contingency . 259,000
- 10% of Direct and Indirect Costs
Fixed Capital Investment 2,978,300
Working Capital 330,900
- 10% of Total Capital Incestment
Total Capital Investment $3,309,200
94
-------
Direct Costs—Process 2B—Continuous, Simultaneous Extraction—Vacuum
Stripped Oil
Purchased Equipment—Installed (End '79) $1,036,400
Instrumentation and Controls—Installed 96,900
- 9.35% of Installed Equipment Costs
Piping—Installed 231,100
- 22.3% of Installed Equipment Costs
Electrical—Installed 74,600
- 7.2% of Installed Equipment Costs
Buildings—Including Services 215,600
- 20.8% of Installed Equipment Costs
Yard Improvements 74,600
- 7.2% of Installed Equipment Costs
Services Facilities—Installed 410,400
- 39.6% of Installed Equipment Costs
Total Direct Plant Cost— 2,139,600
Indirect Costs
Engineering and Supervision , 261,200
- 25.2% of Installed Equipment Costs
Construction Expense , 216,600
- 20.9% of Installed Equipment Costs
Total Direct and Indirect Costs 2,617,400
Contractor's Fee , 130,900
- 5% of Direct and Indirect Costs
Contingency 261,700
- 10% of Direct and Indirect Costs
Fixed Capital Investment 3,010,000
Working Capital 334,500
- 10% of Total Capital Investment
Total Capital Investment $3,344,500
95
-------
Manufacturing and Product Costs
Labor Requirements [16]—
Process 1-A—
Extractor—1st Stage
Extractor—2nd Stage
Evaporator
Vacuum Evaporator
Operating Labor
Men Required Per Shift
1
1
1
1
4 men
3 shifts i 8 hr , $7.00
day
Process 1-B—
shift
Vacuum Evaporator
Extractor—1st Stage
Extractor—2nd Stage
Evaporator
Vacuum Evaporator
hr
345 day
operating year
$231,840
1
1
1
1
1
_ i 3 shifts i 8 hr
j men —5 , . _
1 day ' shift
Process 2-A—
Extractor
Separator
Evaporator
Vacuum Evaporator
$7.00 i 345^day
hr ' operating year
3 1/2 men
3 shifts i 8 hr
day
Process 2-B—
Vacuum Evaporator
Extractor
Separator
Evaporator
Vacuum Evaporator
4 1/2 men
3 shifts
day
shift
8 hr
shift
$7.00
hr
345 day
= $289,800
1
1/2
1
1
operating year
3 1/2
| = $202,860
1
1
1/2
1
1
4 1/2
$7.00 i 345 day
hr ' operating year
= $260,820
96
-------
Raw Materials Cost [16] —
Base Gal/Year $l/Year
Process 1-A—
Pyrolytic Oil
MIBK
Process 1-B —
Pyrolytic Oil
MIBK
Process 2- A —
Pyrolytic Oil
MIBK
Process 2-B —
Pyrolytic Oil
MIBK
Utility Summary —
Basis: 100% Capacity;
Steam —
Demand (#/hr)
Cost $/#(1000)
Cost Per Year ($)
Cooling Water —
Gal/Hr
Gal /da
Cost: $/1000 gal
Cost/yr ($)
Electricity — Estimate
#/day product
kwhr/da
Cost $/kwhr
Cost/yr
.24/gal
.34/#
($2.272/gal)
.24/gal
.34/#
($2,272/gal)
. .24/gal
.34/#
($2.272/gal)
.24/gal
.34/#
($2.272/gal)
345 days /year
7,100,000
111,700
7,100,000
77,400
7,100,000
112,900
7,100,000
67,900
1-A 1-B 2-A
21,540 16
2.30 2
410,200 313
7,640 16
183,460 391
0.07 0
4,430 9
0.10 kwhr/# Product
228,700 215
22,870 21
.042
331,386 311
,440 17,510
.30 2.30
,100 333,500
,320 7,720
,760 185,370
.07 0.07
,460 4,480
[16]
,200 198,700
,520 19,870
042 .042
,825 287,916
1,704,000
253,800
1,957,800
1,704,000
175,800
1,879,800
1,704,000
256,500
1,960,500
1,704,000
154,300
1,858,300
2-B
13,360
2.30
254,400
15,670
376,140
0.07
9,080
190,700
19,070
.042
276,324
97
-------
Process 1-A—
Manufacturing Cost—
Raw Materials 1,957,800
Operating Labor 231,800
Operating Supervision + Clerical
- 15% of Operating Labor 34,800
Utilities
Steam 410,200
Cooling Water 4,400
Electricity 331,400
Maintenance and Repairs
- 7% of Fixed Capital Investment/yr 229,000
Operating Supplies
- 15% of Total Cost of M + R 34,300
Laboratory Charges
- 15% of Operating Labor 34,800
Direct Production Costs 3,268,500
Fixed Charges - (Depreciation, Taxes, Insurance,
Rent) - 10% of Total Product Cost 390,700
Plant Overhead Costs
- 50% of (Operating Labor+Supervision+Maintenance) 247,800
Total Product Cost 3,907,000
Process 1-B—
Manufacturing Cost—
Raw Materials 1,879,800
Operating Labor 289,800
Operating Supervision + Clerical
- 15% of Operating Labor 43,500
Utilities
Steam 313,100
Cooling Water 9,500
Electricity 311,800
Maintenance and Repairs
- 7% of Fixed Capital Investment/yr 238,300
Operating Supplies
. - 15% of Total Cost of M + R 35,700
Laboratory Charges
••;- 15% of Operating Labor 43,500
Direct Production Costs 3,165,000
Fixed Charges - (Depreciation, Taxes, Insurance,
Rent) - 10% of Total Product Cost 383,400
Plant Overhead Costs
- 50% of (Operating Labor+Supervision+Maintenance) 285,800
Total Product Cost 3,834,200
98
-------
Process 2-A—
Manufacturing Cost—
Raw Materials 1,960,500
Operating Labor 202,900
Operating Supervision + Clerical
- 15% of Operating Labor 30,400
Utilities
Steam 333,500
Cooling Water 4,500
Electricity 287,900
Maintenance and Repairs
- 7% of Fixed Capital Investment/yr 208,500
Operating Supplies
- 15% of Total Cost of M + R 31,300
Laboratory Charges
- 15% of Operating Labor ' 30.400
Direct Production Costs 3,089,900
Fixed Charges - (Depreciation, Taxes, Insurance,
Rent) - 10% of Total Product Cost 367,800
Plant Overhead Costs
- 50% of (Operating Labor + Supervision+Maintenance) 220,900
Total Product Cost 3,678,600
Process 2-B—
Manufacturing Cost—
Raw Materials 1,858,300
Operating Labor 260,800
Operating Supervision + Clerical
- 15% of Operating Labor 39,100
Utilities
Steam 254,400
Cooling Water 9,100
Electricity 276,300
*Maintenance and Repairs
- 7% of Fixed Capital Investment/yr 210,700
Operating Supplies
- 15% of Total Cost of M + R 31,600
Laboratory Charges
- 15% of Operating Labor 39.100
Direct Production Costs 2,979,400
Fixed Charges - (Depreciation, Taxes, Insurance,
Rent) - 10% of Total Product Cost 359,400
Plant Overhead Costs
- 50% of (Operating Labor+Supervision+Maintenance 255,300
Total Product Cost 3,594,100
99
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Depreciation
The depreciation over the 10 year life of the plant is shown for each
of the four processes in TABLES 35 to 38. The tables give the annual and
cumulative depreciation based on the double-declining balance method for the
first five years with a switch to straight line depreciation for the remaining
five years. The total depreciable amount includes (installed): equipment,
instrumentation and controls, piping, electrical, buildings and services, yard
improvements, service facilities and land. The total direct plant costs for
each process are shown below:
Process 1A $2,325,768
Process IB 2,419,277
Process 2A 2,117,055
Process 2B 2,139,646
TABLE .35. DEPRECIATION - PROCESS 1A
End of Year
1
2
3
4
5
6
7
8
9
10
Annual
465,154
372,123
297,698
238,159
190,527
152,422
152,422
152,421
152,421
152,421
Cumulative
465,154
837,277
1,134,975
1,373,134
1,563,661
1,716,083
1,868,505
2,020,926
2,173,347
2,325,768
TABLE 36. DEPRECIATION - PROCESS IB
End of Year
1
2
3
4
5
6
7
8
9
10
Annual
483,855
387,084
309,668
247,734
198, 187^
158,550
158,550
158,550
158,550
158,549
Cumulative
483,855
870,939
1,180,607
1,428,341
1,626,528
1,785,078
1,943,628
2,102,178
2,260,728
2,419,277
100
-------
TABLE 37. DEPRECIATION - PROCESS 2A
End of Year
1
2
3
4
5
6
7
8
9
10
Annual
423,411
338,729
270,983
216,786
173,429
138,744
138,744
138,743
138,743
138,743
Cumulative
423,411
762,140
1,033,123
1,249,909
1,423,338
1,562,082
1,700,826
1,839,569
1,978,312
2,117,055
TABLE 38. DEPRECIATION - PROCESS 2B
End of Year Annual Cumulative
1
2
3
4
5
6
7
8
9
10
427,929
342,343
273,875
219,100
175,280
140,224
140,224
140,224
140,224
140,223
427,929
770,272
1,044,147
1,263,247
1,438,527
1,578,751
1,718,975
1,859,199
1,999,423
2,139,646
Products
The products generated by each of the four pyrolytic oil extraction
processes are: Process 1A—insoluble oil, MIBK soluble organics and water
soluble organics; Process IB—volatiles, insoluble oil, MIBK soluble organics
and water soluble organics; Process 2A, water soluble organics and MIBK solu-
ble organics; and Process 2B, volatiles, water soluble organics and MIBK
soluble organics. A survey of the prices of various chemicals was taken with
the results listed in TABLE 39.
The volatiles contain about 68% water with about 20% acetic acid, by
weight. The current market price of acetic acid is $0.23/lb. The selling
price of the volatiles was estimated to be $0.23/lb of acetic acid contained
in the fraction.
101
-------
TABLE 39. PRICE SURVEY OF VARIOUS CHEMICALS [18]
Compound
Alcohol (Synthetic)-(C-12 to C-15)
Acetic Acid
Acetic Anhydride
Acetyldehyde
Acetone
MEK
Ethyl Atnyl Ketone
MIBK
Mineral Spirits
Naptha (VM + P) ,-•
(Varnish + Paint Makers)
Tallow (Fatty Acids-Tech)
Tall Oil (Crude)
Napthol (Tech)
Lacquer Diluent-Pet. Base
Price
$/lb
$/gal
Benzene
Cyclo Hexane
Toluene
Toluene (Coal Tar)
Xylenes
Ortho-Xylene
Para-Xylene
Meta-Xylene
Cumene
Napthalene
Styrene
Para-Tert-Amylphenol
Di-Tert-Amylphenol (85%)
Di-Tert-Amylphenol (95%)
Di-Tert-Amylphenol (97%)
Bis-Phenol-Polycarbonate Grade
Bis-Phenol-Epoxy Grade
Phenol (Synthetic)
Phenyl Acetate
Acetophenone
Benzaldehyde
Benzophenone
Benzyl Alcohol
Bisphenol-A Epoxy Grade
Diphenyl (99.9%)
0-Phenyl Phenol
P-Phenyl Phenol
.225
.279
.17
.19
.185
.22
.28
.31
.15
.25
.35
.74
.61
.78
.79
.61
.57
.38
1.04
.40
1.05-1.15
2.80-2.85
1.00-1.09
.47- .52
.495
1.35-2.00
1.10-1.25
1.65
1.75
1.25
1.35
1.35
.45
.23
.34
.265,
.26
.31
.38
• 34
.43
.32- .49
($150-160/ton)
1.03
.38- .40
.38
(continued)
102
-------
TABLE 3g (continued)
Compound
Price
$/lb
$/gal
Epoxy Resin
Sucrose (#2)
Asphalt
Coal Tar Pitch
Creosote-Coal Tar
Creosote (80/20 Solution)
M-Cresol (95-98%)
M,P-Cresol (90%)
M,P-Cresol (94%)
M,P-Cresol (97%)
0-Cresol (98%)
0-Cresol (99%)
P-Cresol (98%)
Cresylic Acid-Coal Tar Der.
(Resin Grade)
Molasses
.93
.25
.085
.083
.081
.98
.54
.55
.70
.55
.555
1.08
.54
($170/ton)
.55-.65
.83
.81
($26/100#)
The insoluble oil product is a heavy oil, somewhat similar to Bunker C.
It was estimated to be comparable to coal tar pitch ($170 per ton or $0.085
per pound) or creosote - coal tar ($0.83 per gallon or $0.08 per pound). The
insoluble oil fraction was given a selling price of $0.08 per pound in the
minimum selling price case and a selling price of $0.09 per pound in the aver-r
age and maximum selling price cases.
The uses of the water soluble organics and the MIBK soluble organics,
the major products, have been discussed in detail. Some of the possible uses
are, to review: a rubber oil additive, an epoxy intermediate, a resin feed-
stock, and an antioxidant additive for rubber. Prices of similar types of
chemicals are: Styrene - $0.35/lb, Napthalene - $0.25/lb, Acetophenone -
$0.40/lb, Bisphenol A Epoxy Grade - $0.47 to $0.52/lb, Cresylic Acid - $0.54/
Ib, 0-Cresol - $0.55/lb, M-Cresol - $0.98/lb, P-Cresol - $1.08/lb, and Mixed
Cresols - $0.54 to $0.70/lb.
The estimated range of the selling price of the organics was $0.30 to
$0.60/lb. These figures are based on the pounds of organics contained in a
given quantity of product solution. Thus the water soluble organics and the
MIBK soluble organics were given a selling price of $0.30/lb for the minimum
selling price case, $0.50/lb for the average selling price case, and'1 $0.60/lb
for the maximum selling price case.
Sales Income
As shown in the products section, the main products, the MIBK soluble
organics and the water soluble organics are estimated to have a selling price
in the range of $0.30 per pound to $0.60 per pound. The average selling
103
-------
price, estimated by comparing the current market price of similar chemicals,
is $0.50 per pound.
The sales income, in dollars per year, is shown below for 3 cases:
minimum selling price ($0.30/lb), average selling price ($0.50/lb), and maxi-
mum selling price ($0.60/lb). The sales figures are based on a 24 hour per
day operation, 345 day operating year at 100% capacity.
SALES INCOME—MINIMUM SELLING PRICE
Process 1-A—
Insoluble Oil
MIBK Soluble Organics
Water Soluble Organics
Basis
0.08/lb
0.30/lb
0.30/lb
Quantity
Produced
5049.6 Ib/hr
877.3 Ib/hr
3603.4 Ib/hr
$/Yr
3,344,860
2,179,210
8.950,850
$14,474,920
Process 1-B—
Volatiles
Insoluble Oil
MIBK Soluble Organics
Water Soluble Organics
0.23/lb
0.08/lb
0.30/lb
0.30/lb
269.1 Ib/hr
5259.3 Ib/hr
565 Ib/hr
2873.8 Ib/hr
512,470
3,483,760
1,403,460
7,138,530
$12,538,220
Process 2-A—
Water Soluble Organics
MIBK Soluble Organics
Process 2-B—
Volatiles
Water Soluble Organics
MIBK Soluble Organics
SALES INCOME—AVERAGE SELLING PRICE
Basis
0.30/lb
0.30/lb
0.23/lb
0.30/lb
0.30/lb
Quantity
Produced
6186.4 Ib/hr
2094 Ib/hr
269.1 Ib/hr
3954.2 Ib/hr
2722.8 Ib/hr
$/Yr
15,367,020
5.201,500
$20,568,520
512,470
9,822,230
6.763.440
$16,585,670
Process 1-A—
Insoluble Oil
MIBK Soluble Organics
Water Soluble Organics
Basis
0.09/lb
0.50/lb
0.50/lb
104
Quantity
Produced
5049.6 Ib/hr
877.3 Ib/hr
3603.4 Ib/hr
$/Yr
3,762,960
3,632,020
14,918.080
$22,313,060
-------
Process 1-B—
Volatiles
Insoluble Oil
MIBK Soluble Organics
Water Soluble Organics
0.23/lb
0.09/lb
0.50/lb
0.50/lb
269.1 Ib/hr
5259.3 Ib/hr
565 Ib/hr
2873.8 Ib/hr
512,470
3,919,230
2,339,100
11,897,530
$18,668,330
Process 2-A—
Water Soluble Organics
MIBK Soluble Organics
Basis
0.50/lb
0.50/lb
Quantity
Produced
6186.4 Ib/hr
2094 Ib/hr
$/Yr
25,611,700
8,669,160
$34,280,860
Process 2-B—
Volatiles
Water Soluble Organics
MIBK Soluble Organics
0.23/lb
0.50/lb
0.50/lb
269-1 Ib/hr
3954.2 Ib/hr
2722.8 Ib/hr
512,470
16,370,390
11,272,390
28,155,250
SALES INCOME—MAXIMUM SELLING PRICE
Process 1-A—
Insoluble Oil
MIBK Soluble Organics
Water Soluble Organics
Process 1-B—
Volatiles
Insoluble Oil
MIBK Soluble Organics
Water Soluble Organics
Process 2-A—
Water Soluble Organics
MIBK Soluble Organics
Basis
0.09/lb
0.60/lb
0.60/lb
0.23/lb
0.09/lb
0.60/lb
0.60/lb
Basis
0.60/lb
0.60/lb
$/Yr
Quantity
Produced
5049.6 Ib/hr 3,762,960
877.3 Ib/hr 4,358,430
3603.4 Ib/hr 17,901,690
$26,023,080
269-1 Ib/hr
5259.3 Ib/hr
565 Ib/hr
2873.8 Ib/hr
512,470
3,919,230
2,806,920
14,277.040
$21,515,660
Quantity
Produced
6186.4 Ib/hr
2094 Ib/hr
$/Yr
30,734,040
10.403,000
$41,137,040
105
-------
Process 2-B—
Volatiles 0.23/lb 269-1 Ib/hr 512,470
Water Soluble Organics 0.60/lb 3954.2 Ib/hr 19,644,470
MIBK Soluble Organics 0.60/lb 2722.8 Ib/hr 13.526.870
$33,683,810
RATE OF RETURN ANALYSIS
To obtain a rate of return discounted cash flow for the life of the
plant, the following method was adopted: the plant life was assumed to be ten
years beginning at year zero with the total initial investment spread over one
year and ending at year zero.
Although an operating plant would be brought to full capacity gradually
(i.e., 50% capacity 1st year, 75% capacity 2nd year, 100% capacity 3rd year
on), for simplicity of calculation it was assumed that the plant would operate
at 100% capacity over the 10-year period.
It was assumed that the initial investment was the sum of the fixed
capital investment plus working capital. At the end of year 10, salvage was
assumed to be zero, but the working capital would be recovered.
The depreciation schedules were calculated using the double-declining
balance method for the first five years, switching to straight line deprecia-
tion for years 6-10.
Cash flows were calculated for each of the four processes before and
after taxes, and are presented in TABLES 40 through 48. Taxes are 46% of
gross profit. The average annual profit and return on investment (ROI) were
calculated on an after tax basis.
For each process two cases were examined, in which annual sales income
was varied. The change in annual sales income is based on the minimum selling
price of $0.30 per pound and the average selling price of $0.50 per pound of
the soluble organic products. The ROI for each of the cases is shown in
TABLE 40 and expressed as a percent.
TABLE 40. RETURN ON INVESTMENT—SUMMARY
Process Product Selling Price
$0.30/lb $0.50/lb
1A
IB
2A
2B
156.31%
123.60%
274.95%
209.10%
272.74%
211.12%
498.72%
395.90%
106
-------
TABLE 41. CASH FLOW—PROCESS 1-A— Case I—$0.30/lb
Year
0
1
2
3
4
5
6
7
8
9
10
Deprec.
465,154
372,123
397,698
238,159
190,527
152,422
152,422
152,421
152,421
152,421
Cumulative
Cash Position
Before Taxes
(3,635,416)
7,397,598
18,337,581
29,203,139
40,009,158
50,767,545
61,487,827
72,208,109
82,928,390
93,648,671
104,368,952
Gross
Profit -
Dep.
10,102,706
10,195,737
10,270,162
10,329,701
10,377,333
10,415,438
10,415,438
10,415,439
10,415,439
10,415,439
Taxes
4,647,245
4,690,039
4,724,275
4,751,662
4,773,573
4,791,101
4,791,101
4,791,102
4,791,102
4,791,102
Gross
Profit -
Taxes
5,920,615
5,877,821
5,843,585
5,816,198
5,794,287
5,776,759
5,776,759
5,776,758
5,776,758
5,776,758
Net Profit
+
Deprec.
6,385,769
6,249,944
6,141,283
6,054,357
5,984,814
5,929,181
5,929,181
5,929,179
5,929,179
5,929,179
Cash Position
After Taxes
(3,635,416)
2,750,353
9,000,297
15,141,580
21,195,937
27,180,751
33,109,932
39,039,113
44,968,292
50,897,471
56,826,650
Fixed Capital Investment $3,271,874
Total Capital Investment $3,635,416
Sales (for each year) = $14,474,920
Manufacturing cost (for each year) = $3,907,060
Gross Profit (for each year) = $10,567,860
"Average Annual Profit = $56,826,650/10 years = $5,682,665
ROI = $5,682,665/3,635,416 * 100% = 156.31%
-------
TABLE 42. CASH FLOW—PROCESS 1-A—CASE II—$0.50/lb
o
00
Year
0
1
2
3
4
5
'6
7
8
9
10
Deprec .
465,154
372,123
297,698
238,159
190,527
152,422
152/422
152,421
152,421
152,421
Cumulative
Cash Position
Before Taxes
(3,635,416)
15,235,738
34,013,861
52,717,559
71,361,718
89,958,245
108,516,667
127,075,089
145,633,510
164,191,931
182,750,352
Gross
Profit-
Dep.
17,940,846
18,033,877
18,108,302
18,167,841
18,215,473
18,253,578
18,253,578
18,253,579
18,253,579
18,253,579
Taxes
8,252,789
8,295,583
8,329,819
8,357,207
8,379,118
8,396,646
8,396,646
8,396,646
8,396,646
8,396,646
Gross
Prof it -
Taxes
10,153,211
10,110,417
10,076,181
10,048,793
10,026,882
10,009,354
10,009,354
10,009,354
10,009,354
10,009,354
Net Profit
+
Deprec.
10,618,365
10,482,540
10,373,879
10,286,952
10,217,409
10,161,776
10,161,776
10,161,775
10,161,775
10,161,775
Cash Position
After Taxes
(3,635,416)
6,982,949
17,465,488
27,839,368
38,126,320
48,343,729
58,505,505
68,667,281
78,829,056
88,990,831
99,152,605
Fixed Capital Investment $3,271,874
Total Capital Investment $3,635,416
Sales (for each year) =$22,313,060
Manufacturing cost (for each year) = $3,907,060
Gross Profit (for each year) = $18,406,000
Average Annual Profit = $99,152,605/10 years - $9,915,261
ROI = $9,915,261/3,635,416 * 100% - 272.74%
-------
TABLE 43. CASH FLOW—PROCESS 1-B—CASE I->$0.30/lb
o
vo
Year
0
1
2
3
4
5
6
7
8
9
10
Deprec .
483,855
387,084
309,668
247,734
198,187
158,550
158,550
158,550
158,550
158,549
Cumulative Gross '-
Cash Position Prof it -
Before Taxes Dep.
(3,782,363)
5,405,512
14,496,616
23,510,304
32,462,058
41,364,265
50,226,835
59,089,405
67,951,975
76-, 814, 545
85,677,114
Fixed
Total
8,220,165
8,316,936
8,394,352
8,456,286
8,505,833
8,545,470
8,545,470
8,545,470
8,545,470
8,545,471
Capital Investment
Capital Investment
Taxes
3,781,276
3,8*25,791
3,861,402
3,889,892
3,912,683
3,930,916
3,930,916
3,930,916
3,930,916
3,930,916
= $3,404,127
- $3,782,363
Gross
Profit -
Taxes
'' .
4,922,744
4,878,229
4,842,618
4,814,128
4,791,337
4,773,104
4,773,104
4,773,104
4,773,104
4,773,104
Net Profit
+
Deprec.
5,406,599
5,265,313
5,152,286
5,061,862
4,989,524
4,931,654
4,931,654
4,931,654
4,931,654
4,931,653
Cash Position
After Taxes
(3,782,363)
1,624,236
6,889,550
12,041,836
17,103,698
22,093,222
27,024,876
31,956,529
36,888,183
41,819,837
46,751,490
Sales (for each year) = $12,538,220
Manufacturing cost (for each year) = $3,834,200
Gross Profit (for each year) = $8,704,020
Average Annual Profit = $46,751,490/10 years = $4,675,149
ROI = $4,675,149/$3,782,363 * 100% = 123.
-------
TABLE 44. CASH FLOW—PROCESS 1-B—CASE II--$0.50/lb
Year
0
1
2
3
4
5
6
7
8
9
10
Deprec.
483,855
387,084
309,668
247,734
198,187
158,550
158,550
158,550
158,550
158,549
-
Cumulative
Cash Position
Before Taxes
(3,782,363)
11,535,622
26,756,836
41,900,634
56,982,498
72,014,815
87,007,495
102,000,175
116,992,855
131,985,535
146,978,214
Gross
Profit -
Dep.
14,350,275
14,447,046
14,524,462
14,586,396
14,635,943
14,675,580
14,675,580
14,675,580
14,675,580
14,675*581
-- • •• •• -
Taxes
6,601,127
6,645,641
6,681,253
6,709,742
6,732,534
6,750,767
6,750,767
6,750,767
6,750,767
6,750,767
Gross
Profit -
Taxes
8,233,004
8,188,489
8,152,877
8,124,388
8,101,596
8,083,363
8,083,363
8,083,363
8,083,363
8,083,363
Net Profit
+
Deprec .
8,716,859
8,575,573
8,462,545
8,372,122
8,299,783
8,241,913
8,241,913
8,241,913
8,241,913
8,241,912
Cash Position
After Taxes
(3,782,363)
4,934,496
13,510,068
21,972,614
30,344,736
38,644,519
46,886,432
55,128,345
63,370,258
71,612,172
79,854,085
Fixed Capital Investment = $3,404,127
Total Capital Investment = $3,782,363
Sales (for each year) = $18,668,330
Manufacturing cost (for each year) =$3,834,200
Gross Profit (for each year) = $14,834,130
Average Annual Profit = $79,854,085/10 years = $7,985,409
ROI = $7,985,409/$3,782,363 * 100% = 211.12%
-------
TABLE 45. CASH FLOW-PROCESS 2-A—CASE I—$0.30/lb
Year
0
1
2
3
4
5
6
7
8
9
10
Deprec.
423,411
338,729
270,983
216,786
173,429
138,744
138,744
138,743
138,743
138,743
Cumulative
Cash Position
Before Taxes
(3,309,177)
14,004,134
31,232,763
48,393,646
65,500,332
82,563,661
99,592,305
116,620,949
133,649,592
150,678,235
167,706,878
Gross
Profit -
Dep.
16,466,489
16,551,171
16,618,917
16,673,114
16,716,471
16,751,156
16,751,156
16,751,157
16,751,157
16,751,157
Taxes
7,574,585
7,613,539
7,644,702
7,669,632
7,689,577
7,705,532
7,705,532
7,705,532
7,705,532
7,705,532
Gross
Profit -
Taxes
9,315,315
9,276,361
9,245,198
9,220,268
9,200,323
9,184,368
9,184,368
9,184,368
9,184,368
9,184,368
Net Profit
+
Deprec.
9,738,726
9,615,090
9,516,181
9,437,054
9,373,752
9,323,112
9,323,112
9,323,111
9,323,111
9,323,111
Cash Position
After Taxes
(3,309,177)
6,429,549
16,044,639
25,560,821
34,997,874
44,371,626
53,694,739
63,017,851
72,340,962
81,664,073
90,987,185
Fixed Capital Investment = $2,978,259
Total Capital Investment = $3,309,177
Sales (for each year) = $20,568,520
Manufacturing cost (for each year) = $3,678,620
Gross Profit (for each year) = $16,889,900
Average Annual Profit = $90,987,185/10 years = $9,098,719
ROI = $9,098,719/$3,309,177 * 100% = 274.95%
-------
TABLE 46. CASH FLOW—PROCESS 2-A—CASE II—$0.50/lb
Year
0
1
2
3
4
5
6
7
8
9
10
'
Deprec.
423,411
338,729
270,983
216,786
173,429
138,744
138,744
138,743
138,743
138,743
Cumulative Gross
Cash Position Profit -
Before Taxes Dep.
(3,309,177)
27,716,474
58,657,443
89,530,666
120,349,692
151,125,361
181,866,345
212,607,329
243,348,312
274,089,295
304,830,278
Fixed
Total
30,178,829
30,263,511
30,331,257
30,385,454
30,428,811
30,463,496
30,463,496
30,463,497
30,463,497
30,463,497
Capital Investment
Capital Investment
Gross
Profit -
Taxes Taxes
13,882,261 16,719,979
13,921,215 16,681,025
13,952,378 16,649,862
13,977,309 16,624,931
13,997,253 16,604,987
14,013,208 16,589,032
14,013,208 16,589,032
14,013,208 16,589,032
14,013,208 16,589,032
14,013,208 16,589,032
= $2,978,259
•- $3,309,177
Net Profit
+
Deprec.
17,143,390
17,019,754
16,920,845
16,841,717
16,778,416
16,727,776
16,727,776
16,727,775
16,727,775
16,727,775
Cash Position
After Taxes
(3,309,177)
13,834,213
30,853,967
47,774,811
64,616,529
81,394,944
98,122,720
114,850,496
131,578,271
148,306,046
165,033,821
Sales (for each year) = $34,280,860
Manufacturing cost (for each year) = $3,678,620
Gross Profit (for each year) = $30,602,240
Average Annual Profit = $165,033,821/10 years = $16,503,382
ROI = $16,503,382/$3,309,177 * 100% = $498,72%
-------
TABLE 47. CASH FLOW—PROCESS 2-B--CASE I—$0.30/lb
u>
Year
0
1
2
3
4
5
6
7
8
9
10
Deprec.
427,929
342,343
273,875
219,100
175,280
140,224
140,224
140,224
140,224
140,223
Cumulative
Cash Position
Before Taxes
(3,344,489)
10,074,940
23,408,783
36,674,158
49,884,758
63,051,538
76,183,262
89,314,986
102,446,710
115,578,434
128,710,157
Gross
. Profit -
Dep.
12,563,571
12,649,157
12,717,625
12,772,400
12,816,220
12,851,276
12,851,276
12,851,276
12,851,276
12,851,277
Taxes
5,779,243
5,818,612
5,850,108
5,875,304
5,895,461
5,911,587
5,911,587
5,911,587
5,911,587
5,911,587
Gross
Profit -
Taxes
7,212,257
7,172,888
7,141,393
7,116,196
7,096,039
7,079,913
7,079,913
7.079,913
7.079,913
7.079,913
Net Profit
+
Deprec.
7,640,186
7,515,231
7,415,268
7,335,296
7,271,319
7,220,137
7,220,137
7,220,137
7,220,137
7,220,136
Cash Position
After Taxes
(3,344,489)
4,295,697
11,810,928
19,226,196
26,561,492
33,832,810
41,052,947
48,273,085
55,493,222
62,713,359
69,933,495
Fixed Capital Investment = $3,010,040
Total Capital Investment <= $3,344,489
Sales (for each year) = $16,585,670
Manufacturing cost (for each year) = $3,594,170
Gross Profit (for each year) = $12,991,500
Average Annual Profit = $69,933,495/10 years = $6,993,350
ROI = $6,993,350/$3,344,489 * 100% = 209.10%
-------
TABLE 48. CASH FLOW—PROCESS 2-B—CASE II—$0.50/lb
Year
0
1
2
3
4
5
6
7
8
9
10
Deprec.
427,929
342,343
273,875
219,100
175,280
140,224
140,224
140,224
140,224
140,223
Cumulative
Cash Position
Before Taxes
(3,344,489)
21,644,520
46,547,943
71,382,898
96,163,078
120,899,438
145,600,742
170,302,046
195,003,350
219,704,654
244,405,958
Gross
Profit -
Dep.
24,133,151
24,218,737
24,287,205
24,341,980
24,385,800
24,420,856
24,420,856
24,420,856
24,420,856
24,420,857
Taxes
11,101,249
11,140,619
11,172,114
11,197,311
11,217,468
11,233,594
11,233,594
11,233,594
11,233,594
11,233,594
Gross
Profit -
Taxes
13,459,831
13,420,461
13,388,966
13,363,769
13,343,612
13,327,486
13,327,486
13,327,486
13,327,486
13,327,486
Net Profit
+
Deprec.
13,887,760
13,762,804
13,662,841
13,582,869
13,518,892
13,467,710
13,467,710
13,467,710
13,467,710
13,467,709
Cash Position
After Taxes
(3,344,489)
10,543,271
24,306,075
37,968,916
51,551,785
65,070,677
78,538,387
92,006,097
105,473,807
118,941,518
132,409,927
Fixed Capital Investment = $3,010,040
Total Capital Investment = $3,344,489
Sales (for each year) = $28,155,250
Manufacturing cost (for each year) = $3,594,170
Gross Profit (for each year) = $24,561,080
Average Annual Profit = $132,409,927/10 years = $13,240,993
ROI = $13,240,993/$3,344,489 * 100% = 395.90%
-------
Since the ROI for each case presented in TABLE 40 is extremely high, and
not normally encountered in practice, a second method was adopted to determine
the profitability of the pyrolytic oil separation processes. Using all the
assumptions previously stated, three rates of return were selected - 15%, 30%,
and 50%, and the selling price of the products necessary to produce this ROI
was calculated. For this analysis the product streams were totaled and all
products were assumed to have the same selling price per pound. The results
are presented in TABLE 49. The rates of return are on an after tax basis. The
required average selling price per pound of product varies from $0.0543 per
pound to $0.1063 per pound. This corresponds to a raw pyrolytic oil cost
(feedstock) of $0.24 per gallon, based on $2.30 per MM BTU, or $0.023 per
pound.
TABLE 49. MINIMUM SELLING PRICE PER POUND TO JUSTIFY INVESTMENT
Process Average Annual Price Per Pound of Product to Total Product
Profit Required Generate the Given ROI 106 Ib/yr
15% 30% 50%
1A $ 545,300 0.0543 78.91
1,090,600 0.0671
1,817,700 0.0842
IB 567,400 0.0569 74.23
1,134,700 0.0711
1,891,200 0.0899
2A 496,400 0.0587 68.56
992,800 0.0721
1,654,600 0.0899
2B 501,700 0.0686 57.51
1,003,300 0.0847
1,672,200 0.1063
115
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SECTION 8
DISCUSSION
PYROLYSIS OF LIGNOCELLULOSIC MATERIALS AND PROPERTIES OF THE OILS
The pyrolysis of lignocellulosic and similar materials produces char,
organic substances, water and gases. Condensation of the off-gas stream from
the pyrolysis will yield an organic phase and an aqueous phase. Pyrolysis of
pine sawdust on a batch basis and the products has been described in detail
[11]. The oils produced from pyrolytic processes were the focus of this
investigation with the emphasis on characterization and maximum resource-
recovery by processing to produce more useful fractions for chemical appli-
cations.
Pyrolytic oils contain a wide spectrum of organic compounds, both
aromatic and aliphatic. Most of these compounds are oxygenated, and cpnse-
quently, the oils contain many functional groups. The oils must be consid-
ered therefore as a chemical feedstock and as a source of chemical materials
for industrial applications. In order to develop the potential of pyrolytic
oils as a chemical feedstock, characterization of chemical and physical
properties of the pyrolytic oils is absolutely necessary. The data from the
characterization of the oils can then be used in the development of processes
for the oils to yield fractions that have chemical applications or that can
be further refined or processed to yield useful chemical products.
The production of pyrolytic oils is an important and significant fac-
tor in the overall utilization of these oils. Some of the factors that
affect the quality and characteristics of the oils are feed materials,
pyrolysis mode (vertical bed reactor, flash pyrolysis, fluidized bed reactor,
etc.), pyrolysis conditions (temperature, presence or absence of air, feed
material size, etc.) and recovery mode from the off-gas stream. For this
investigation, the oils were obtained from the Tech-Air Corporation's 50 dry
ton/day pyrolysis facility and the pyrolysis pilot plants on the Georgia Tech
campus which utilize the Georgia Tech process [6, 7]. Oils obtained from
pilot plants and field demonstration units which operate on a continuous
basis at steady state conditions are representative of the oils that would
be produced on a commercial scale. Therefore, the results and data obtained
from a study with these oils will be more applicable in the processing and
recovery of useful products from commercially produced pyrolytic oils.
Pyrolytic oils from the Tech-Air 50 dry ton/day facility have been
thoroughly characterized as to overall general properties such as heating
value, elemental content, acidity, etc., and these results have been reported
116
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[12]. In general, the oils are dark brown to black and have a pungent, burnt
odor. The viscosity of the oils will depend upon the amount of water present
in the oils. The water is well emulsified and does not separate on standing.
With a water content of 25% or greater the oils are relatively thin and free
flowing. Oils which are essentially free of water are viscous, and some have
a grease-like consistency at ambient temperature. The viscosity decreases
with temperature for short periods of heating. The oils are heat sensitive
and prolonged heating will result in increasing viscosity with eventual for-
mation of solids. The oils are combustible and can be burned very satisfac-
torily with the proper burner. Burning tests of the oils admixed with fuel
oil and with char have been very satisfactory. The oils are acidic and exhi-
bit corrosive properties. This characteristic must be taken into account in
the storage and processing of the oils.
f\ ">
ANALYSIS-1AND CHARACTERIZATION OF PYROLYTIC OILS
The pyrolytic oils are a complex mixture of organic compounds with a
wider range in boiling point from highly volatile substances to very high
boiling substances. The oils contain oxygen in the range of 20 to 40 percent,
and therefore, there is a large number of oxygen containing compounds present
which have a variety of organic functional groups such as carbonyl, hydroxyl,
ether, etc. The oils are heat sensitive and begin to decompose at 175° to
200°C. The chemical and physical analysis of pyrolytic oils therefore is not
a simple task. It is very difficult to analyze the oils, or fractions obtained
frotethe oils, for chemical content as to classes of organic compounds and as
to functionality. This aspect of this project has been very difficult, and
there is a need for additional work in the chemical analysis of pyrolytic oils.
!. • ''
=fe ' r For overall properties of pyrolytic oils, many of the ASTM methods are
applicable to pyrolytic oils. The tests used in characterizing and analyzing
pyrolytic oils from the Tech-Air Cordele development unit are given in TABLE 2.
The distillation test, ASTMD-86, is not too useful as most of the oils start
to decompose at the point when approximately 50 to 60 percent of the oil has
distilled. The development of more meaningful tests will be necessary as pyro-
lytic oils find greater utility as fuels.
..-(.:>:; .•*...• The chemical composition of the pyrolytic oils are of importance and
significance in developing processing methods for the oils. A knowledge of
thercomponents in the oils will serve as guidelines for devising processing
methods for separation of the oils into fractions containing a major chemical
class of substances, i.e., phenolics. The oils are chemically complex and
contain a wide variety of aliphatic, aromatic and heterocyclic compounds. The
analytical techniques that are very useful and valuable in determining the
composition of the oils are liquid chromatography (LC), gas chromatography
(GC), thin layer chromatography (TLC), gas chromatography/mass spectroscopy
(GC/MS), and infrared spectroscopy. .
Considerable effort was placed on both liquid and gas chromatography in
this investigation and both were used extensively in this work with the oils.
LC is an excellent analytical technique for these oils as it is carried out at
ambient temperature,'is capable of high resolution of complex organic mixtures
and component detection is nondestructive. The oils are heat sensitive, and
117
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hence, LC is particularly useful with these oils. In addition, the pyrolytic
oils are soluble in organic-water solvent systems which are very useful in LC.
A variety of LC columns and conditions were tested in determining suitable
conditions for analyzing pyrolytic oils. Particular interest was in using LC
as a "finger-printing" method for the oils and fractions obtained from them
by various processing techniques. A Partisil ODS 5p column with a water-
acetonitrile solvent system and a flow rate of one ml per minute was found to
produce very satisfactory chromatograms. In Phase III of the experimental
work, the column used for LC was a 25 cm Spherosal ODS CIQ column. The most
useful ultraviolet detector settings are 280 nm and 254 nm for our purposes.
LC was used throughout this investigation for the "finger-printing" of the
oils and fractions of the oils obtaining by different processing methods.
Gas chromatography (GC) offers an excellent technique for analyzing
complex mixture of organic compounds. The disadvantage with GC with pyrolytic
oils is the heat sensitivity of the oils since GC analysis involves tempera-
tures up to 250°C for these oils. Recognizing this as a possible constraint:,
GC should be useful for analysis of fractions containing more volatile compo-
nents, of water soluble components and of fractions obtained in experiments
designed to separate the raw oils into fractions containing a major chemical
class of compounds. A variety of column packings and conditions were tested.
A column containing 10% methylsilicone fluid has been found to be very useful
with the raw oils and fractions with higher boiling components, and a column
containing 10% Carbowax 20M has been found satisfactory for low boiling com-
ponents. In Phase III of the experimental work, a Pora Pak Q column was used
for water and water soluble organics.
Thin layer chromatography was utilized in Phase III of the experimental
program as it offered a very rapid and useful technique for analyzing the dif-
ferent phases and fractions obtained in the extraction experiments. Details
are given in Phase III of the experimental section.
A nonaqueous titration method was devised and used to determine the
presence of phenolics in the oils and fractions obtained from the oils by the
various extraction techniques. The technique has utility with thesefoils rso
long as the limitations are recognized. More details are given in the exper-
imental section, Phase III.
In our'attempts to analyze the pyrolytic oils, and particularly the
fractions obtained from the oils by the various processing techniques, the
data from LC, TLC, IR, GC and nonaqueous titration were used and evaluated.
The most promising avenues for improving the chemical analytical data with
the fractions obtained from the oils are to correlate the components obtained
in GC, LC and TLC with GC/MS and IR data. For pilot work with pyrolytic oils,
there is a need for rapid analytical techniques to follow the process during
actual operation. TLC may offer a potential method for this need.
DISTILLATION
The distillation of complex liquids is a widely used process that has
reached a high degree of sophistication in the chemical industry. Therefore,
'distillation offers a possible method for processing and refining pyrolytic
oils. It is particularly useful for obtaining fractions with fairly close
118
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boiling range. A number of distillation experiments were conducted with the
raw oils. These included distillation at atmospheric pressure, and at
0.2 - 0.4 mm mercury, fractional distillation at reduced pressure, steam dis-
tillation and vacuum stripping.
The distillation of the raw oils at both atmospheric and low pressures
would yield from 55 to 65 percent distillate. The charge in the flask would
become more viscous as the distillation proceeded, and when 55 to 65 percent
had distilled, the remaining oil in the flask would begin to decompose and
smoke. In some cases, the charge would decompose quickly with an evolution
of gases. From these experiments, it was concluded that distillation of the
raw oils was not a suitable first step for processing the oils.
The distillate from a simple vacuum distillation of raw pyrolytic oil
was^fractionated at approximately 2 mm pressure. The distillation did not
yield any fractions with a close boiling range. The liquid chromatograms of
the fractions indicated, however, that the more polar and water soluble com-
ponents were concentrated in the low boiling fractions whereas the less polar
components were concentrated in the higher boiling fractions. A sample of
water-insoluble oil, which had been prepared by the water extraction of vacuum
stripped oil, was distilled at approximately 6 mm pressure. Approximately,
47 percent of the sample distilled from 50°C up to 193°C, and no close-boiling
fractions were obtained. Analysis of the five fractions by TLC, LC, and IR
indicated that the first four fractions, approximately 29 percent of the,
charge, contained mainly phenolic aromatics and phenolic ethers. The chemical
analyses indicated that fraction 5, boiling point range 175° - 193°C and
approximately 17 percent of the charge, contained mainly phenolic ethers and
aromatic neutrals with a trace of polyhydroxy neutral compounds. The results
with the distillation of fractions obtained from raw pyrolytic oil samples ,
show that distillation of oil fractions produced from raw oil by other separa-
tion techniques can be used to yield more highly refined chemical materials.
F j .Steam distillation of pyrolytic oil samples showed that a relatively
small amount of the oils were steam distilled. The steam distillate contained
more polar and water soluble components of the oil. The liquid chromatogram
of, the steam distillate was very similar to the liquid chromatograms of the
water extracts at 25°C, 50°C and 95°C of the oil, indicating that steam dis-
tillation and water extraction of the oils produced very similar fractions of
the oils. These results indicate that steam distillation is not a suitable
processing step for the raw oils. Vacuum stripping of the oils .,at ambient
temperature was found to be an effective way to remove the water and some of
the volatile organics, which included the acids.
HYDROGENATION
Hydrogenation was considered as a possible means of improving the pro-
cessing characteristics of the oils. A series of hydrogenation experiments
with raw pyrolytic oil samples were carried out at 4 atmospheres and 20 atmos-
pheres. The Pd catalyst performed better than the Pt catalyst. The amount of
hydrogen absorbed was in the range of 2 mg/g of oil from pine wood and 3 to 5
mg/g of oil from hardwood. If one assumes an average molecular weight of 150
for the oil, then approximately 0.15 mole of hydrogen is absorbed.per mole of
oil. for 2 mg of hydrogen absorbed/g of oil. These preliminary experiments at
119
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relative low pressure Indicated that hydrogenation should not be considered
as an initial step in processing pyrolytic oils. Additional hydrogenation
experiments should be conducted with fractions of oil obtained by various
separation techniques including some at higher pressures than used in this
work.
EXTRACTION EXPERIMENTS
Some factors that are important and must be considered in developing
processing technology to produce fractions of pyrolytic oil that are suitable
for chemical applications are the wide spectrum of organic compounds present
in the oils, the quantity of each compound is relatively low, the oils are
heat sensitive and chemically reactive, the solubility characteristics of the
oils, and the volatiles (boiling point 100°C or less) including water in the
raw oils. Two chemical operations that seemed most appropriate to investigate
as processing steps were distillation and extraction. Distillation has been
discussed above, and based on our results fractional distillation at reduced
pressure on oil fractions obtained by an extraction process should be seriously
considered as an operation to yield highly refined products. The focus of
the processing study and the major effort was with extraction methods. The
study initially was based on bench scale experiments with five different
approaches. Based on the results of these experiments, three processes were
selected for further investigation with batch processing. For the continuous
extraction experiments, two processes were selected from the batch processing
studies for investigation with both raw pyrolytic oil and vacuum stripped oil.
Bench Scale Extraction Experiments
Five major approaches involving extraction techniques were tested at
the bench level. These approaches, which are discussed in the experimental
section with results, were as follows:
A - Extraction of oil sequentially with water at 25°C, 50°C and 95°C.
B - Extraction of oil with sodium sulfate solution (salting out effect).
C - Extraction of oil simultaneously with an organic solvent and water
(three phase system).
D - Extraction of sodium hydroxide solution at different pH ranges with
methylene chloride.
E - Extraction of organic solvent solutions of pyrolytic oil with water.
Each of these approaches, or combinations, offer possibilities that can
be utilized in a final process that will result in the production of fractions
of oil for chemical applications. Based on some initial results with both raw
and vacuum stripped oil, it was decided to use vacuum stripped oil in these
batch experiments at the bench level. Treatment of the raw oil at reduced
pressure and ambient temperature removes volatiles (largely acidic) and most
of the water. The significant results for each approach are presented.
From process A, approximately 50% of the original raw oil was isolated
as a water insoluble organic fraction, which contained about 20% phenolics and
80% aromatic neutrals. The separation of this fraction into the two major com-
pound classes is very desirable. Subsequent processing techniques that are
120
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potentially useful are fractional distillation and other extraction steps.
The *3w?ee aqueous fractions, if combined, would contain approximately 34% of
the original oil with 27% phenolics and 73% polyhydroxy neutral substances.
A potential means for separation of these two chemical classes is the use of
the salting-out technique which is the basis of process B. The advantage of
process A is that water is a cheap solvent and nonhazardous, and the process
should be relatively simple.
Process B, which involves essentially a salting out effect with sodium
'sulfate, offers a possibility for separation of the polyhydroxy neutral sub-
stances. The first step would be as depicted in Figure 28. The aqueous
fraction contained organics with approximately 70% phenolics. This separation
could possibly be improved by determining optimum conditions. The insoluble
fraction could be extracted with water to remove the polyhydroxy neutral sub-
stances leaving an insoluble oil fraction. The salting out technique has the
disadvantages of the organics having to be recovered from the concentrated
salt solution and of the recovery and recycling the salt solution.
:Process C, the three phase system, offers some interesting separation
possibilities. It should be noted that the phenolics in the oil are separated
about 50-50 in processes A and B and in process C, Figure 29, the aqueous
phase"contains about 50% of the phenolics and the remaining 50% is about
evenly-divided between the ether phase and the insoluble oil phase. Fractional
distillation of the separate oil and ether fractions should yield fractions
with>high concentration of phenolics and the aromatic neutrals. The three
phase approach with anisole produced results as shown in Figure 30. The
quantities of the components in the aqueous fraction are about the same as
when diisopropyl ether was used. The anisole, however, dissolves a much
greater portion of the oil than diisopropyl ether.
In process D, two percent sodium hydroxide solution was used as a sol-
vent for the oil followed by extraction with methylene chloride at three dif-
ferent pH ranges, 8 to 10, 5 to 7, and 1 to 3. Approximately, 53% of the
oil"charge dissolved in 300 ml of 2% NaOH. The extraction with CH2C12 at
p"Hc8 to'J10 gave predominantly aromatic neutrals whereas at the low pH range,
the extract contained predominantly phenolics. Approximately 55% of the
phenolics were in the aqueous phase with the remainder distributed in the
three CHoCl2 extracts. The remainder of the charge dissolved in 400 ml of
2% NdOH, and the solution was extracted in the same manner as above. It
should be noted that in the first CH^C^ extract, approximately 92% of the
organics was aromatic neutral compounds. Also, in the aqueous phase, approx-
imately 58% of the organics was phenolics. Additional bench scale work is
needed with this process to determine its usefulness as a method of processing
pyrolysis oil. This-approach has the disadvantages that it involves a number
of'processing steps and no one extraction produced a clear fraction of a given
class of compounds present in the pyrolytic oil.
In process E, the organic solvents tested were methylene chloride and
n-bufcanol. Two solutions of pyrolysis oil in methylene chloride were extracted
with water followed by extraction of the aqueous solution in one experiment
with diisopropyl ether and in the second experiment with methyl isobutyl
ketone (MIBK). The results-are shown schematically for the two experiments
121
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in Figures 32 and 33, respectively. A significant result of these two experi-
ments is that the polyhydroxy neutral substances are concentrated in the
aqueous phase along with 50 to 60% of the phenolics in the oil. The methylene
chloride fraction contains phenolics and aromatic neutral compounds which
could be fractional distilled to provide more desirable and useful fractions
of the oil. Another approach to the treatment of the aqueous fraction is
extraction with MIBK. The extraction of the aqueous fraction with MIBK gave
a solution with approximately 92% phenolics, which represents approximately 35%
of the phenolics in the aqueous fraction.
A solution of pyrolysis oil in n-butanol was extracted with water to
determine the separation that would be obtained and the results are shown
schematically in Figure 34. The polyhydroxy neutral substances are distri-
buted between the aqueous fraction and the n-butanol fraction which is not a
desirable result. Consequently, this approach was not pursued.
: An examination of the data from the process approaches discussed above
shows that for each approach approximately 50% of the phenolic content of the
oil is in the aqueous fraction with the remainder in the insoluble oil phase
or in the organic solvent phase. This could be of significance in that each
of these phenolic fractions could have greater utility for specific uses than
a single combined fraction of the phenolics. The aqueous fractions from all
of the approaches contain relatively large amounts of polyhydroxy neutral
substances with the exception of the salting out techniques.
The aqueous insoluble fractions contain approximately 50% of the phe-
nolic content of the oil along with most of the aromatic neutral compounds
with ratios of phenolics to aromatic neutral compounds in the range of 1 to 3
and 1 to 4. The separation of this fraction into the two major classes of
compounds could possibly be accomplished by fractional distillation or
extraction with an alkaline solution.
Careful examination of the data from the bench scale experiments with
the five processes and consideration of each overall process as a continuous
chemical process indicated that processes A and C are the most promising with
process E offering some potential.
Continuous Extraction Experiments
The extraction experiments and related work for this phase of the
program is described in the Experimental Section*, Phase III. The pyrolytic
oil used in these experiments was produced in the Georgia Tech pyrolysis pilot
plant under carefully controlled conditions in October, 1978, from pine cljips.
Based on the results from the bench scale extraction experiments
described above, the decision was made to investigate further the three
extraction methods, listed below, with both raw and vacuum stripped pyrolytic
oil.
•Process No. 1 - Water Extraction
•Process No. 2 - Simultaneous Extraction with Water and an Organic
Solvent
•Process No. 3 - Dissolution in an Organic Solvent Followed by
Water Extraction
122
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Additional batch experiments were conducted with 100, 200 or 500 g oil
samples using the three approaches. The results with Process No. 1 in this
phase were comparable to the results in Experimental Phase II with aqueous
extraction. The results indicate that MIBK is a better solvent than chloro-
form for extraction of the aqueous phase. Based on the results and observa-
tions of the experiments with Process No. 2, MIBK was selected as the solvent
for the continuous extraction experiments. Also, in these batch experiments
with the three phase system, the insoluble tar phase was very small, 2 percent
or less, whereas in the initial experiments with the three phase process using
diisopropyl ether, the insoluble phase was 25 percent. With Process No. 3, an
insoluble tar phase was present in each experiment. It was decided to dis-
continue experimentation with this approach as it did not appear to offer any
advantage over the simultaneous use of water and an organic solvent, Process
N6.:; 2.
The continuous countercurrent experiments were conducted with Process
No. 1 and Process No. 2 with both raw and vacuum stripped pyrolytic oil. Some
Important observations from the results of the continuous extraction experi-
ments are that the polyhydroxy neutrals are essentially concentrated in the
aqueous phases f6r all four experiments, that the aromatic neutrals in the
aqueous phase are extracted essentially completely into the MIBK fraction
along with some phenolics, and that the insoluble oil phases of Process No. 1
and the MIBK phases of Process No. 2 contain phenolics and aromatic neutrals.
The MIBK phases and fractions and the insoluble oil phases which contain
mainly phenolics and aromatic neutrals could be further processed by frac-
tional distillation. Concentration of the extracted aqueous fractions, which
contain phenolics (approximately 15 to 30 percent) and polyhydroxy neutrals,
from both processes could yield a solution from which additional phenolics
could be extracted. The results of these experiments are very promising that
pyrolytic oils can be processed by extraction techniques to yield fractions
that have potential for chemical applications or that can be refined through
additional chemical processing operations.
iiJlV XJj, ' '
PILOT''MANT : '
Based on the data obtained from the continuous countercurrent extrac-
tion experiments at the bench level, a versatile pilot plant was designed which
can be used to test the water extraction Process No. 1 arid the simultaneous
extraction Process No. 2 with water and an organic solvent. The processes can
be tested with both raw oil and vacuum stripped oil at a rate of four gallons
pef^minute. In addition to the various extraction operations, oil fractions
• could be further processed by distillation. The data from the continuous
extraction experiments indicate that the1 extraction approach'is a very promis-
ing one by which fractions of the oil can be obtained which can be processed
by additional operations, particularly fractional distillation, to yield
products of greater utility. With a pilot plant, the concept can be demon-
strated and sufficient quantities of oil fractions can be obtained for testing
and development for industrial applications. More details on the pilot plant
and schematics are given in Section 6.
123
-------
ECONOMICS
In order to make some preliminary economic assessments for processing
pyrolytic oils into materials for chemical applications, it was necessary to
base the analysis on the data from the bench scale countercurrent continuous
extraction experiments with the two processing modes with raw and vacuum
stripped oils. The major assumptions were that pyrolytic oils, either raw or
vacuum stripped, could be processed on a continuous basis by the two proces-
sing modes and that the processing modes would yield oil fractions which would
be suitable for commercial applications. The major objective of this analysis
was to determine if the processing of pyrolytic oils appeared to be economically
feasible.
It was assumed the pyrolytic oil plant would process oil produced by
five wood pyrolysis plants, each processing five dry tons per hour for 345
days per year. The yield of oil was assumed to be 18 percent on a dry weight
basis, which amounts to approximately 7,100,000 gallons per year. The oil
was assumed to have a heating value of 10,000 Btu/lb and a density of 10 Ib
per gallon. The cost of the oil to the plant was based on a value of $2.30
per million Btu.
The analysis was approached in two ways. In one method, the average
selling price per pound for the total output from each process mode was deter-
mined to provide a net return on investment (ROI) for 15, 30, and 50 percent.
The average selling price per pound for each process for this approach is
given in TABLE 50. The significance of this analysis is that it shows that
the selling price — 8.4 to 10.6 cents per pound — for the oil fractions for
a 50 percent return on investment is in the range of quoted market prices in
December, 1979, for similar materials, such as coal tar creosote at 9 cents
per pound and well below the quoted prices for coal tar cresylic acid at 54
cents per pound. In the other method, the analysis was made on the basis that
for case one, the oil product selling prices would be 8 cents per pound for
the insoluble oil, 23 cents per pound for the organic volatiles from the oil
stripping and 30 cents per pound for both the MIBK soluble and water soluble
organics. In case two, the insoluble oil was 9 cents per pound, the organic
volatile fraction, 23 cents per pound and both the MIBK soluble and water
soluble organics, 50 cents per pound. The return on investment for each case
is presented in TABLE 51. Each specific process for each case provides an
TABLE 50. AVERAGE SELLING PRICE FOR PYROLYTIC OIL PRODUCTS
Net Return on Investment
Process
15%
30%
50%
1A
IB
2A
2B
5.4c/lb
5.7
-------
TABLE 51. RETURN ON INVESTMENT - PERCENT
Process Case 1 Case 2
1A
IB
2A
2B
156
124
275
209
273
211
499
396
excellent return on investment. The significance of these results is that the
economic feasibility appears to be very promising for processing the oil into
products for commercialization. In order to realize the potential for proces-
sing pyrolytic oil into chemical materials on a commercial scale, it would be
necessary to study and obtain more data by processing pyrolytic oils with a
small scale pilot plant (see Section 6) and to investigate commercial appli-
cations for oil fractions produced with the pilot plant. In this way, reli-
able operating costs could be established and commercial value of the products
could be determined.
125
-------
REFERENCES
1, E. Hagglund. Chemistry of Wood. Academic Press, Inc., New York, 1952.
2. A. J. Hamm and E. E. Harris. Chemical Processing of Wood. Chemical
Publishing Co., Inc., New York, 1953.
3. L. F. Hawley and L. E. Wise. The Chemistry of Wood. Chemical Catalog
Co., New York, 1926.
4. R. H. Farmer. Chemistry in the Utilization of Wood. Pergamon Press,
Oxford, 1967.
5. L. A. Hawley. Wood Distillation. The Chemical Catalog Co., Inc., New
York, 1923.
6. M. D. Bowen, E. D. Smyly, J. A. Knight, and K. R. Purdy. A Vertical Bed
Pyrolysis System in Solid Wastes and Residues: Conversion by Advanced
Thermal Processes. J."L. Jones and S. B. Radding, eds., pp. 94-125-
ACS Symposium Series 76, American Chemical Society, 1978.
7. J. A. Knight. Pyrolysis of Wood Residues with a Vertical Bed Reactor in
Progress in Biomass Conversion, Volume 1. K." V. Sarkanen. and D. A.
Tillman, eds., pp. 87-115. Academic Press, Inc., 1979.
8. F. L. Rissinger. Changing Feedstocks-Chemicals or Calories? - Chemical
Engineering Progress, 71: 30-33 (1975).
9. L. L. Anderson. Energy Potential from Organic Wastes. U.S. Department
of the Interior. Circular 8549 (1972).
10. D. A. Tillman. Wood as an Energy Source. Academic Press, Inc., New
York, 1978.
11. J. A. Knight. Pyrolysis of Pine Sawdust in Thermal Uses and Properties
of Carbohydrates and Lignins. F. Shafizadeh, K. V. Sarkanen and D. A.
Tillman, eds., pp. 159-173. Academic Press, Inc., 1976.
12. J. A. Knight, D. R. Hurst, and L. W. Elston. Wood Oil from Pyrolysis of
Pine Bark-Sawdust Mixture in Fuels and Energy from Renewable Resources.
D. A. Tillman, K. V. Sarkanen and Larry L. Anderson, eds., pp. 169-195.
Academic Press, Inc., 1977.
126
-------
13. M. B. Polk. Development of Methods for the Stabilization of Pyrolytic
Oils. Annual Report. June, 1977. Grant No. R 804 440 010. U. S.
Environmental Protection Agency, Cincinnati, Ohio 45268.
14. Peters, M. S., and K. D. Timmerhouse. Plant Design and Economics for
Chemical Engineers, 2nd ed., McGraw-Hill Book Co., New York, N. Y., 1968.
15. Perry, R. H., and C. H. Chilton. Chemical Engineers Handbook, 5th ed.,
McGraw-Hill Book Co., New York, N. Y., 1973, Sections 13, 15, 21, 23, '
25.
16. Aries, R. S., and R. D. Newton. Chemical Engineering Cost Estimation.
McGraw-Hill Book Co., New York, N. Y., 1955, pp. 118-182.
17. Guthrie, K. M. Capital Cost Estimating. In: Modern Cost-Engineering
Techniques, H. Popper, ed., McGraw-Hill Book Co., New York, N. Y.,
1970, pp. 80-108.
18. Current Prices of Chemicals and Related Materials. Chemical Marketing
Reporter, V216 (#24): 38-48, Dec. 10, 1979.
'..'
19. The Chemical Rubber Co. Handbook of Chemistry and Physics, 47th ed.,
R. C. Weast, ed. Chemical Rubber Publishing Co., Cleveland, Ohio, 1966,
Sections C, D.
20. Dean, J. A. Lange's Handbook of Chemistry Method. McGraw-Hill Book Co.,
New York, N. Y., 1974, pp. 9-85—9-96.
21. Drew, J. W. Design for Solvent-,Recovery. Chemical Engineering Progress
V71 (No. 2): 92-99, Feb. 1975.
22. McCabe, W. L., and J. C. Smith. Unit Operations of Chemical Engineering,
2nd ed. McGraw-Hill Book Co., New York, N. Y., 1967, pp. 299-321.
23. Combustion Engineering, Inc. Steam Tables—Properties of Saturated and
Superheated Steam, 2nd printing, Windsor, Connecticut, 1967, 35 pp.
24. Mission Analysis for the Federal Fuels From Biomass Program. Volume IV:
Thermochemical Conversion of Biomass to Fuels and Chemicals. Report:
Jan. 1979. SRI International; Menco Park, CA, pp. 105-121.
127
-------
APPENDIX A
MATERIAL BALANCE CALCULATIONS
LABORATORY SCALE—CONTINUOUS EXTRACTION
Process 1-A—Raw Oil-Two Stage Extraction—Total Reactant and Product Balance
Extractor—1st Stage
Raw Oil - 1998g
Nonvolatile Organics - 1698g
Volatile Organics - 140g
Water - 160g
;
Water - 4120g
Aqueous Phase - 4806g
Nonvolatile Organics - 808g
Volatile Organics - 108.7g
Water - 3889.3g
Insoluble Oil Phase - 1121g
Nonvolatile Organics-818g
Volatile Organics - 25.3g
Water - 211.1%
Apparent Losses - 191g
Nonvolatile Organics - 72g
Volatile Organics - 6.0g
Water - 113g
128
-------
TABLE A-l. FLOWRATES—RUN NO. 1-A
HP-29C — Linear Curve Fit*
The data is fitted to a straight line (linear regression) . The form of the
equation is shown below, with x = time, (min.)» y = accumulated stream input
or output (grams) .
y = a + bx
Raw Oil Input Rate
n = 21
Water Input Rate
n = 13
Aqueous Phase Output Rate
n = 21
Insoluble Oil Output Rate
n = 9
a = 10.949
b = 10.5964 grams/min
r2 = 0.99732
a = 8.0936
b = 22.3411 grams/min
r2 - 0.9985
a = -4.785531
b = 26.3342 grams/min
r2 - 0.9983
a = -105.51
b = 5.831 grams/min
r2 - 0.9604
Hewlett-Packard HP-19C-29C Applications Book, p. 102-106.
129
-------
TABLE A-2. SEPARATION PROCESS 1-A—RAW OIL--2 STAGE EXTRACTION--
CUMULATIVE REACTANT AND PRODUCT WEIGHTS
Time
Min.
0
5
10
20
30
50
60
65
70
80
90
100
110
120
130
140
150
160
170
180
185
Input
Raw Oil*
g
0
61
126
237
361
534
607
667
760
906
1001
1092
1165
1231
1328
1446
1643
1741
1789
1935
1998
Water
g
0
120
250
470
570
1000
1190
1360
1570
1830
2000
2250
2460
2630
2880
3150
3420
3620
3820
4000
4120
Output
Aqueous Phase
g
0
115
292
594
813
1209
1439
1605
1861
2158
2449
2679
2960
3127
3336
3732
3982
4170
4482
4785
4805
Oil
g
0
-
-
-
89
-
200
226
-
-
369
-
-
512
-
-
738
-
-
927
1121
Raw Oil Density = 1.215 g/ml
Process 1-B—Vacuum Stripped Oil—Two Stage Extraction—Total Reactant and
Product Balance
• — '• i "•'• • • — — i .. \
Vacuum Evaporator-(Stripper)—
Raw Oil - 1932g
Water - 226.3g
Organics - 1705.7g
Steam
Volatiles - 284g
Water - 226.3g
Organics - 57.7g
167°F Vacuum Stripped Oil - 1648g
Nonvolatile Organics-1621g
Volatile Organics - 27g
75°C
130
-------
The raw pyrolytic oil had a water content of 11.712% and an organics
content of 88.288%. The purpose of the stripping operation was to remove the
water from the pyrolytic oil. However, small scale tests to strip the oil
showed that the raw pyrolytic oil had a volatiles content of 14.7% (composed
of 79.57% water and 20.43% organics). Thus, some organics had been volatilized
in the process of stripping the water from the pyrolytic oil.
Extractor-lst Stage—
Vacuum Stripped Oil-1648g
Nonvolatile Organics-1621g
Volatile Organics - 27g
Water - 2830g
Aqueous Phase - 3194g
Nonvolatile Organics - 667g
Volatile Organics - 19.3g
Water - 2507.7g
Insoluble Oil Phase-1129g
Nonvolatile Organics-903g
Volatile Organics - 6.8g
Water - 219.2g
Apparent Losses - 155g
Nonvolatile Organics - 51g
Volatile Organics - 0.9g
Water - 103.Ig
TABLE A-3. FLOWRATES—RUN NO. IB
HP-29C—Linear Curve Fit* ,,
The data is fitted to a straight line (linear regression). The form of the
equation is
y = a + bx
where x = time (min); y = accumulated stream input or output (grams)
Vacuum Stripped Oil Rate
n = 7
Water Input Rate
n = 10
Aqueous Phase Output Rate
n = 10
Insoluble Oil Output Rate
n = 4
a - -19.259
b = 13.098 grams/min
r2 = 0.9741
a = -34.406
b = 23.297 grams/min
r2 = 0.9711
a = -58.849
b = 25.617 grams/min
r2 = 0.9925
a = -34.49
b = 8.777 grams/min
r2 = 0-9837
Hewlett-Packard HP-19C/HP-29C Applications Book, p. 102-106.
131
-------
TABLE A-4. SEPARATION PROCESS 1-B—VACUUM STRIPPED OIL—2 STAGE
EXTRACTION—CUMULATIVE REACTANT AND PRODUCT WEIGHTS
Input
Output
Time
Mln.
0
10
20
50
60
70
75
90
115
130
Vacuum
Stripped Oil*
g
0
42
105
336
672
787
1113
1217
1501
1648
Water
g
0
200
490
1000
1150
1400
2000
2280
2750
2830
Aqueous
Phase
g
0
190
502
1054
1476
1606
1949
2310
3013
3194
Oil
g
0
-
-
-
402
—
657
-
-
1129
Vacuum stripped oil density = 1.238 g/ml
Process 2-A—Raw Oil—Simultaneous MIBK and Water Extraction—Total Reactant
and Product Balance
Extractor—
Raw Oil - 555.3g
Nonvolatile Organics-472g
Volatile Organics - 38.9g
Water - 44.4g
MIBK - 562g
Water - 780g
Aqueous Phase - 1153g
Nonvolatile Organics-358g
Volatile Organics - 23.6g
Water - 748.3g
MIBK - 23.1 g
MIBK Phase - 741g
Nonvolatile Organics-114g
Volatile Organics-15.2g
Water - 74.2g
MIBK - 537.6g
Apparent Losses - 3.3g
Nonvolatile Organics - Og
Volatile Organics - O.lg
Water - 1.9g
MIBK - 1.3g
132
-------
TABLE A-5. FLOWRATES—RUN No. 2-A
HHBH^VIBBIIBHBIIBBBHIBBVBB
HP-29C—Linear Curve Fit*
The data is fitted to a straight line (linear regression). The form of the
equation is
y = a + bx
where x = time (min); y = accumulated stream input or output (grams)
Raw Oil Input Rate a = 19.5851
n - 7 b = 6.19224
r2 = 0.98756
Water Input Rate a = -31.19097
n = 7 b = 9.38692
r2 = 0.98394
MIBK Input Rate a = -28.9572
n = 7 b = 6.17643
r2 = 0.98622
Total Output Rate a = -74.80375
21.8058]
0.99509
n = 7 b = 21.80581
2
Hewlett-Packard HP-19C/HP-29C Applications Book, p. 102-106.
TABLE A-6. SEPARATION PROCESS 2-A—RAW OIL—SIMULTANEOUS EXTRACTION WITH
MIBK AND WATER—CUMULATIVE REACTANT AND PRODUCT WEIGHTS
Time
Min.
0
10
22
30
53
75
90
Raw Oil*
g
0
49.4
129.6
197.4
377.6
499.7
555.3
Input
Water
g
0
70
150
190
500
720
780
MIBK
g
0
24.1
96.5
160.8
265.3
418.0
562.0
Output
Total
g
0
140
380
534
1009
1601
1918
Raw Oil Density = 1.234 g/ml
133
-------
Process 2-B—Vacuum Stripped Oil—Simultaneous MIBK and Water Extraction-
Total Reactant and Product Balance
Vacuum Evaporator-(Strlpper)—
Raw Oil - 1967.2g
Water - 230.4g
Organics - 1736.8g
Steam
Volatiles - 289.2g
Water - 230.4g
Organics - 58.8g
Vacuum Stripped Oil-1678g
Nonvolatile Organics-1629g
Volatile Organics - 49g
Extractor—
Vacuum Stripped Oil-1678g _
Nonvolatile Organics - 1629g
Volatile Organics - 49g
Water - 1900g
MIBK - 1198g
Aqueous Phase - 2746g
Nonvolatile Organics-735g
Volatile Organics - 31.4g
Water - 1790.Ig
MIBK - 189.5g
MIBK Phase - 1812g
Nonvolatile Organics-798g
Volatile Organics-15.7g
Water - 36.2g
MIBK - 962:lg
Apparent Losses - 218g
Nonvolatile Organics-96g
Volatile Organics -1.9g
Water - 73.7g
MIBK - 46.4g
All inputs were measured quantities, as were the quantities in the
Aqueous Phase and the MIBK Phase. The total amount of Apparent Losses was
found by difference. Nonvolatile Organic content of the Aqueous Phase and
the MIBK Phase was measured. Nonvolatile Organic content in Apparent Losses
was determined by difference.
The remaining constituents of Apparent Losses (solvents and volatile
organics) were calculated as explained below. Losses occurred by two methods,
spillage and evaporation. It was assumed that the losses due to spillage
were much greater than the losses due to evaporation. Thus, the losses of
134
-------
solvents and volatile organics of "apparent losses" will occur in the same
proportion as their proportion in the well-mixed extractor fluid.
The percentage of volatile organics in each of the output streams was
estimated to be the same as the percent volatile organics in the solvents
and volatile organics portion of the input stream.
The four remaining components were the amounts of water and MIBK in
both the aqueous and the MIBK phases. As stated previously the extractor
effluent was a well mixed dispersion. The effluent was allowed to stand
overnight to separate into 2 phases. But even after overnight settling,
some MIBK remained dissolved and/or mixed in the Aqueous Phase and some
water remained dissolved and/or unseparated in the MIBK phase. For design
purposes it was estimated that the water content of the MIBK phase was 2%.
The remainder of the mass balance was calculated. The resulting MIBK content
of the aqueous phase was 6.9%.
TABLE A-7. FLOWRATES—RUN NO. 2-B
HP-29C Linear Curve Fit*
The data is fitted to a straight line (linear regression). The form of the
equation is
y = a + bx
where x = time (min) ; y = accumulated stream input or outp.ut, (grams).
Vacuum Stripped Oil Rate a = -45.7056
0 b = 13.2809
n = o «
r = 0.99105
Water Input Rate a = 1.55937
b = 15.540
n = 12 7
r = 0.9944
MIBK Input Rate a = 48.94185
b = 9.74807
n = 11 9
r = 0.97695
Total Output Rate a = -197.6378
! b =''38.4628
n = 9 r2 = 0.961935
*Hewlett-Packard HP-19C/HP-29C Applications Book, p. 102-106.
135
-------
TABLE A-8. SEPARATION PROCESS 2-B--VACUUM STRIPPED OIL—SIMULTANEOUS
EXTRACTION WITH MIBK AND WATER—CUMULATIVE REACTANT
AND PRODUCT WEIGHTS
Time
Min.
0
10
30
45
60
70
80
90
100
115
120
125
Input
Vacuum Stripped Oil
g
0
30
70
347
595
864
983
1092
1231
1469
1614
1678
Water
g
0
160
480
670
920
1000
1320
1460
1620
1800
1820
1900
MIBK
g
0
24
281
458
723
804
892
973
1045
1138
1164
1198
Output
Total
g
0
213
Ilk
1113
1858
2159
2551
3650
4116
4255
4306
4558
*
Vacuum Stripped Oil Density = 1.238g/ml
136
-------
APPENDIX B
PILOT PLANT CALCULATIONS
MAJOR EQUIPMENT—MATERIAL BALANCES
Process 1-A—Raw Oil—Two Stage Extraction
Extractor—1st Stage
Raw Oil - 2432.6 Ib/hr
Nonvolatile Organics -
2067.3 Ib/hr
Volatile Organics
170.5 Ib/hr
Water - 194.8 Ib/hr
Water - 5016.2 Ib/hr
Aqueous Phase-6083.9
Nonvolatile Organics
1071.4 Ib/hr
Volatile Organics -
139.7 Ib/hr
Water - 4872.8 Ib/hr
Insoluble Oil Phase
1364.8 Ib/hr
Nonvolatile Organics
995.9 Ib/hr
Volatile Organics-30
Water - 338.1 Ib/hr
Ib/hr
.8 Ib/hr
Raw Oil = 4 gal/min
ft'
62.4 Ib
ft'
1.215
60 min
hr
= 2432.6 Ib/hr
7.48 gal
Nonvolatile Organics = 2432.6 Ib/hr |1698/1998| = 2067.3 Ib/hr
Volatile Organics = 2432.6 Ib/hr |140/1998] = 170.5 Ib/hr
Water - 2432.6 Ib/hr |160/1998| - 194.8 Ib/hr
Water = 2432.6 Ib/hr |4120/1998| - 5016.2 Ib/hr - 602 gal/hr
Aqueous Phase = 2432.6 Ib/hr |(4806+191)/1998| = 6083.9 Ib/hr
137
-------
Nonvolatile Organics = 6083.9 Ib/hr |(808+72)/499?| = 1071.4 Ib/hr
Volatile Organics = 6083.9 Ib/hr |(108.7 + 6.0)/4997| = 139.7 Ib/hr
Water = 6083.9 Ib/hr |(3889.3 + 113)/4997| = 4872.8 Ib/hr
Insoluble Oil Phase = 2432.6 Ib/hr |1121/1998) = 1364.8 Ib/hr
Nonvolatile Organics - 1364.8 Ib/hr |818/1121| = 995.9 Ib/hr
Volatile Organics - 1364.8 Ib/hr (25.3/1121| =30.8 Ib/hr
Water - 1364.8 Ib/hr |277.7/112l| = 338.1 Ib/hr
Extractor-2nd Stage
Aqueous Phase-6083.9 Ib/hr
Nonvolatile Organics -
1071.4 Ib/hr
Volatile Organics -
139.7 Ib/hr
Water - 4872.8 Ib/hr
MIBK - 2436.6 Ib/hr
r
1200
2400
ml MIBK
g AQ PHASE
MIBK Soluble Phase -
2673.7 Ib/hr
MIBK - 2436.6 Ib/hr
Organics - 237.1 Ib/hr
Water Soluble Phase -
5846.8 Ib/hr
Water - 4872.8 Ib/hr
r\-v cr an •{ r^ey Q 7 "^ Q"7 1 "K /Tit*
urganics jijtji j-o/nr
O.SOlg „,., , 1K/K-.-
, Z'tju.O J.D/111.
ml
MIBK = 6083.9 Ib/hr
From laboratory analysis 19.58% of the organics in the Aqueous Phase input
stream were present in the MIBK soluble phase, and 80.42% of the organics in
the Aqueous Phase input stream were present in the Water Soluble Phase.
MIBK Soluble Phase
Organics = 1211.1 Ib/hr |.1958| = 237.13 Ib/hr
Water Soluble Phase
Organics - 1211.1 Ib/hr |.8042J - 973.97 Ib/hr
138
-------
Evaporator—
MIBK Soluble Phase - 70°F
2673.7 Ib/hr
MIBK - 2436.6 Ib/hr
Organics-237.1 Ib/hr
Steam - 358.43°F
150 psia saturated
249°F MIBK - 2436.6 Ib/hr
244°F Organics-237.1 Ib/hr
244°F Condensate-432 Ib/hr
MIBK = 2436.6 Ib/hr
(244 - 70)°F| +
2436.6 Ib/hr |82.5 BTU/lb| + 2436.6 Ib/hr | °5^ ™ \
(249 - 244°F)| = 401,213 BTU/hr
BTU
Organics'- (Estimate cp to be 0.55
Ib
) =
237.1 Ib/hr
0.55 BTU
Ib • °F
(244 - 70)°F = 22,690 BTU/hr
Total = 423,903 BTU/hr
Steam Use = x Ib/hr |(1194.1 BTU/lb - 1162.0 BTU/lb)| +
x Ib/hr |949.5 BTU/lb| = 423,903 BTU/hr
x = 432 Ib/hr, 150 psia sat steam
s
Vacuum Evaporator—
Water Soluble Phase -
5846.8 Ib/hr
Water - 4872.8 Ib/hr
Organics - 973.97 Ib/hr
70°F
Steam - 5390 Ib/hr
150 psia saturated
358.43°F
170°F Water - 4872.8 Ib/hr
220°F Organics-973.97 Ib/hr
139
-------
Water = 4872.8 Ib/hr |(1134.2 BTU/lb - 137.97 BTU/lb|
+ 4872.8 Ib/hr |TT~?= | (170 - 70°F| - 5,341,710 BTU/hr
Organics = 973.97 Ib/hr \
Total = 5,422,062 BTU/hr
(220 - 70)°F| = 80,353 BTU/hr
Steam Use = x Ib/hr | (1194.1 BTU/lb - 1153.4 BTU/lb) |
+ x Ib/hr | 965. 2 BTU/lb | = 5,422,062 BTU/hr
x = 5390 Ib/hr steam, 150 psia saturated
Process 1-B — Vacuum Stripped Oil — Two Stage Extraction
Vacuum Evaporator- (Stripper) —
Raw Oil - 2432.6 Ib/hr
Water - 284.9 Ib/hr
Organics - 2147.7 Ib/hr
Steam - 499 Ib/hr
70°F
358.43°F
170°F Volatiles - 357.6 Ib/hr
Water - 284.9 Ib/hr
Organics - 72.7 Ib/hr
220°F Vacuum Stripped Oil -
2075.0 Ib/hr
Nonvolatile Organics -
L 2041.0 Ib/hr
Volatile Organics-34.0 ib/hr
Raw Oil = 4 gal/min
ft"
62.4 Ib
7.48 gal ' ft3
!.215
Water = 2432.6 Ib/hr |.117l| = 284.9 Ib/hr
Organics = 2432.6 Ib/hr |.8829| = 2147.7 Ib/hr
Volatiles - 2432.6 Ib/hr |.147| = 357.6 Ib/hr
Organics = 357.6 Ib/hr |.2032J = 72.7 Ib/hr
Water = 357.6 Ib/hr |.7968| = 284.9 Ib/hr
Vacuum Stripped Oil = 2432.6 Ib/hr |.853| - 2075.0 Ib/hr
Nonvolatile Organic = 2075.0 Ib/hr |.9836| = 2041.0 Ib/hr
Volatile Organic - 2075.0 Ib/hr |.0164| = 34.0 Ib/hr
140
-------
Volatiles
Water - 284.9 Ib/hr |j^°.B°f 1(170 - 70)°F| + 284.9 Ib/hr J996.2 BTU/lb|
= 312,307 BTU/hr
Organics = 72.7 Ib/hr I^i5] LBTU I (170 - 70)°F | + 72.7 Ib/hr |l95.5 BTU/lb|
= 17,969 BTU/hr
Vacuum Stripped Oil - 2075.0 Ib/hr \^5.
= 171,188 BTU/hr
Total » 501,464 BTU/hr
(220 - 70)°F|
Steam Use - x Ib/hr | (1194.1 BTU/lb - 1153.4 BTU/lb) | + x Ib/hr
•|965.2 BTU/lb| - 501,464 BTU/hr
x = 499 Ib/hr steam, 150 psia saturated
Extractor-lst Stage' —
Vacuum Stripped Oil -
2075.0 Ib/hr
Nonvolatile Organics -
2041.0 Ib/hr
Volatile Organics -
34.0 Ib/hr
Water - 3563.3 Ib/hr
Aqueous Phase-4216.7 Ib/hr
Nonvolatile Organics -
904.0 Ib/hr
Volatile Organics -
25.4 Ib/hr
Water - 3287.3 Ib/hr
Insoluble Oil Phase -
1421.6 Ib/hr
Nonvolatile Organics -
1137.0 Ib/hr
Volatile Organics-8.6 Ib/hr
Water - 276.0 Ib/hr
Water = 2075.0 Ib/hr |2830/1648] = 3563.3 Ib/hr
Aqueous Phase - 2075.0 Ib/hr |(3194 + 155)/1648| = 4216.7 Ib/hr
Nonvolatile Organics = 4216.7 Ib/hr |(667 + 51)/3349| = 904.0 Ib/hr
Volatile Organics - 4216.7 Ib/hr |(19.3 + 0.9)/3349| = 25.4 Ib/hr
Water • 4216.7 Ib/hr |(2507.7 + 103.1)/3349| = 3287.3 Ib/hr
141
-------
Insoluble Oil Phase = 2075.0 Ib/hr |1129/1648[ = 1421.6 Ib/hr
Nonvolatile Organics = 1421.6 Ib/hr |903/1129| = 1137.0 Ib/hr
Volatile Organics = 1421.6 Ib/hr |6.8/1129) =8.6 Ib/hr
Water = 1421.6 Ib/hr |219.2/1129| = 276.0 Ib/hr
Extractor-2nd Stage
Aqueous Phase - 4216.7 Ib/hr
Nonvolatile Organics -
904.0 Ib/hr
Volatile Organics -
25.4 Ib/hr
Water - 3287.3 Ib/hr
MIBK - 1688.8 Ib/hr
MIBK = 4216.7 Ib/hr
MIBK Soluble Phase -
1841.5-Ib/hr
MIBK - 1688.8 Ib/hr
Organics - 152.7 Ib/hr
Water Soluble Phase -
4064.0 Ib/hr
Water - 3287.3 Ib/hr
Organics - 776.7 Ib/hr
1200 ml MIBK
2400g Aq Phase
0.801E
ml
= 1688.8 Ib/hr
From laboratory analysis 16.43% of the organics in the Aqueous Phase input
stream were present in the MIBK soluble phase, and 83.57% of the organics
in' the aqueous phase input stream were present in the water soluble phase.
MIBK Soluble Phase
Organics = 929.4 Ib/hr .1643 = 152.7 Ib/hr
Water Soluble Phase
Organics = 929.4 Ib/hr .8357 = 776.7 Ib/hr
Evaporator—
MIBK Soluble Phase - 70°F
1841.5 Ib/hr
MIBK - 1688.8 Ib/hr
Organics - 152.7 Ib/hr
Steam 358.43°F
249°F MIBK - 1688.8 Ib/hr
244°F Organics - 152.7 Ib/hr
244°F Condensate - 298 Ib/hr
142
-------
MIBK = 1688.8 Ib/hr
[(244 - 70)°F| +
1688.8 Ib/hr J82.5 BTU/lb| + 1688.8 Ib/hr
I 0.459 BTUi
1 Ib • °F '
(249-244)°F| = 278,080 BTU/hr
' BTU
Organics - (Estimate cp to be
) = 152.7 Ib/hr
|0>55 BTU
•| (244 - 70)°F | = 14,613 BTU/hr
Total = 292,693 BTU/hr
'Steam Use = x Ib/hr ((1194.1 BTU/lb - 1162.0 BTU/lb)|
+ x Ib/hr [949-5 BTU/lb = 292,693 BTU/hr
x = 298 Ib/hr Steam, 150 psia sat
Vacuum Evaporator
Water Soluble Phase -
4064.0 Ib/hr
Water - 3287.3 Ib/hr
Organics - 776.7 Ib/hr
70°F
Steam - 3646 Ib/hr
150 psi saturated
348.43°F
170°F Water - 3287.3 Ib/hr
220°F Organics - 776.7 Ib/hr
Water - 3287.3 Ib/hr |(1134.2 BTU/lb - 137.97 BTU/lb)|
i1.0 BTU,
3287.3 Ib/hr
Organics = 776.7 Ib/hr
(170 - 70)°F = 3,603,637 BTU/hr
|(220 - 70)°F| = 64,078 BTU/hr
Total = 3,667,715 BTU/hr
Steam Use = x Ib/hr | (1194.1 BTU/lb - 1153.4 BTU/lb) | +
x Ib/hr | 965. 2 BTU/lb | = 3,667,715 BTU/hr
x = 3,646 Ib/hr steam, 150 psia saturated
143
-------
Process 2-A—Raw Oil—Simultaneous MIBK and Water Extraction
Extractor
Raw Oil - 2470.6 Ib/hr
Nonvolatile Organics -
2100.0 Ib/hr
Volatile Organics -
173.1 Ib/hr
Water - 197.5 Ib/hr
MIBK - 2500.4 Ib/hr
Water - 3470.3 Ib/hr
Raw Oil
8al/min
ft
3 62.4
Overhead Effluent -
8441.3 Ib/hr
Nonvolatile Organics -
2100.0 Ib/hr
Volatile Organics -
173.1 Ib/hr
Water - 3667.8 Ib/hr
MIBK - 2500.4 Ib/hr
Insoluble Oil Phase - 0 Ib/hr
.4 . i 9Q/. 1 f.n m-Sn /V>»- — O/,7n £ 1U/V,v
7.48 gal
ft"
Nonvolatile Organics = 2470.6 Ib/hr J472/555.3J = 2100.0 Ib/hr
Volatile Organics = 2470.6 Ib/hr 138.9/555.31 = 173.1 Ib/hr
Water = 2470.6 Ib/hr j44.4/555.3| = 197.5 Ib/hr
MIBK = 2470.6 Ib/hr )562/555.3| = 2500.4 Ib/hr
Water = 2470.6 Ib/hr |780/555.3| = 3470.3 Ib/hr
Separator
Overhead Effluent -
8441.3 Ib/hr
Nonvolatile Organics -
2100.0 Ib/hr
Volatile Organics -
' 173.1 Ib/hr
Water - 3667.8 Ib/hr
MIBK - 2500.4 Ib/hr
Aqueous Phase - 5144.5 Ib/hr
Nonvolatile Organics -
1592.8 Ib/hr
Volatile Organics -
105.4 Ib/hr
Water - 3337.7 Ib/hr
MIBK - 10.8.6 Ib/hr
MIBK Phase - 3296.8 Ib/hr
Nonvolatile Organics -
507.2 Ib/hr
Volatile Organics-67.7 Ib/hr
Water - 330.1 Ib/hr
MIBK - 2391.8 Ib/hr
144
-------
Aqueous Phase = 8441.3 Ib/hr |(1153 + 3.3)/1897.3| = 5144.5 Ib/hr
Nonvolatile Organics = 5144.5 Ib/hr |358/(1153 + 3.3)| = 1592.8 Ib/hr
Volatile Organics = 5144.5 Ib/hr |(23.6 + O.I)/(1153 + 3.3)| = 105.4 Ib/hr
Water = 5144.5 Ib/hr |(748.3 + 1.9)7(1153 + 3.3)| = 3337.7 Ib/hr
MIBK = 5144.5 Ib/hr |(23.1 + 1.3)/(1153 + 3.3)| = 108.6 Ib/hr
MIBK Phase = 8441.3 Ib/hr |741/1897.3| = 3296.8 Ib/hr
Nonvolatile Organics = 3296.8 Ib/hr |114/741| = 507.2 Ib/hr
Volatile Organics = 3296.8 Ib/hr |15.2/741| =67.7 Ib/hr
Water = 3296.8 Ib/hr J74.2/74l| = 330.1 Ib/hr
MIBK = 3296.8 Ib/hr |537.6/741| = 2391.8 Ib/hr
Evaporator (or Column)—
MIBK Soluble Phase -
3296.8 Ib/hr
Nonvolatile Organics -
507.2 Ib/hr
Volatile Organics -
67.7 Ib/hr
Water - 330.1 Ib/hr
MIBK - 2391.8 Ib/hr
Steam
150 psia saturated
MIBK = 2391.8 Ib/hr
70°F
r
r
r
358.43°F
249°F MIBK - 2391.8 Ib/hr
244°F Organics - 574.9 Ib/hr
244°F Condensate - 831 Ib/hr
0.459 BTU
Ib • °F
• |82.5 BTU/lb| + 2391.8 Ib/hr
= 393,836 BTU/hr
(244 - 70)°F| + 2391.8 Ib/hr
(249 -
BTU
..,, Organics - (Estimate cp to be 0.55 | ^— OF
• 1(244 - 70)°F| = 55,018 BTU/hr
) =574.9 Ib/hr |
0.55BTUi
Ib • °F '
145
-------
Water = 330.1 Ib/hr
(212 - 70)°FJ + | 330.1 lb/hr|970.3 BTU/lb
= 367,170 BTU/hr
Total = 816,024 BTU/hr
Steam Use = x Ib/hr | (1194.1 BTU/lb - 1162.0 BTU/lb) | + x Ib/hr
•| 949. 5 BTU/lb | = 816,024 BTU/hr
x = 831 Ib/hr steam, 150 psia saturated
Vacuum Evaporator (Double Effect) —
Water Soluble Phase -
5144.5 Ib/hr
Nonvolatile Organics -
1592.8 Ib/hr
Volatile Organics -
105.4 Ib/hr
Water - 3337.1 Ib/hr
MIBK - 108.6 Ib/hr
Steam
150 psia saturated
Water = 3337.7 Ib/hr
70°F
r
r
r
358.43°F
249°F MIBK - 108.6 Ib/hr
244°F Organics - 1698.2 Ib/hr
244°F Condensate - 3963 Ib/hr
(212 - 70)°F| +
82.5 BTU
lb
3337.7 Ib/hr 1970.3 BTU/lb] = 3,712,523 BTU/hr
MIBK - 108.6 Ib/hr [°^4590pTU| (244 - 70)°F + 108.6 Ib/hr
+ 108.6 Ib/hr ^90pTU | (249-244) 9F | - 17,882 BTU/hr
Organics - (estimate cp to be 0.55 ——=Wr) =
lb • F
1698.2 Ib/hr 1°^ o™ |(244 - 70)°F| = 162,518 BTU/hr
Total = 3,892,923 BTU/hr
Steam Use = x Ib/hr ((1194.1 BTU/lb - 1162.0 BTU/lb)|
+ x Ib/hr (949.5 BTU/lb| = 3,892,923 BTU/hr
x = 3966 Ib/hr steam, 150 psia saturated
146
-------
Process 2-B—Vacuum Stripped Oil—Simultaneous MIBK and Water Extraction
Vacuum Evaporator-(Stripper)—
Raw Oil - 2478.6 Ib/hr
Organics - 2188.4 Ib/hr
Water - 290.2 Ib/hr
70°F
Steam
358.43°F
Volatiles 364.4 Ib/hr
Organics - 74.2 Ib/hr
Water - 290.2 Ib/hr
Vacuum Stripped Oil -
Raw Oil = 4 gal/min [ ? ^ ^ \
2114.2 Ib/hr
Nonvolatile Organics -
2052.5 Ib/hr
Volatile Organics-61.7 Ib/hr
1.238 | 60 min/hr| =2478.6 Ib/hr
Water = 2478.6 Ib/hr |.117l| = 290.2 Ib/hr
Organics = 2478.6 Ib/hr |.8829| = 2188.4 Ib/hr
t ,^ Volatiles = 2478.6 Ib/hr |.147| = 364.4 Ib/hr
Organics = 364.4 Ib/hr |.2032| =74.2 Ib/hr
Water = 364.4 Ib/hr |.7968J = 290.3 Ib/hr
Vacuum Stripped Oil = 2478.6 Ib/hr |.853| = 2114.2 Ib/hr
Nonvolatile Organics = 2114.2 Ib/hr |1629/1678] = 2052.5 Ib/hr
Volatile Organics = 2114.2 Ib/hr |49/1678) =61.7 Ib/hr
Volatiles
1.0 BTU i(17Q _ 70)oF| + 290>2 lb/hr |996-2 BTU/lb|
Water = 290.2 Ib/hr
Ib • °F
= 318,217 BTU/hr
Organics = 74.2 Ib/hr
= 18,340 BTU/hr
Vacuum Stripped Oil = 2114.2 Ib/hr
Total = 510,979 BTU/hr
(170 - 70)°F|+ 74.2 Ib/hr |l95.5 BTU/lb|
\ (220 - 70)°F| = 174,422 BTU/hr
Steam Use = x Ib/hr | (1194.1 BTU/lb - 1153.4 BTU/lb) |
+ x Ib/hr | 965. 2 BTU/lb | = 510,979 BTU/hr
x = 508 Ib/hr steam, 150 psia saturated
147
-------
Extractor—
Vacuum Stripped Oil -
2114.2 Ib/hr
Nonvolatile Organics -
2052.5 Ib/hr
Volatile Organics-61.7 Ib/ht
Water - 2393.9 Ib/hr
MIBK - 1509.4 Ib/hr
Overhead Effluent -
6017.5 Ib/hr
Nonvolatile Organics -
2052.5 Ib/hr
Volatile Organics-61.7 Ib/hr
Water - 2393.9 Ib/hr
MIBK - 1509.4 Ib/hr
Insoluble Oil Phase- 0 Ib/hr
MIBK = 2114.2 Ib/hr |1198/1678| - 1509.4 Ib/hr
Water = 2114.2 Ib/hr |l900/1678| = 2393.9 Ib/hr
Separator—
Overhead Effluent -
6017.5 Ib/hr "
Nonvolatile Organics -
2052.5 Ib/hr
Volatile Organics-61.7 Ib/hr
Water - 2393.9 Ib/hr
MIBK - 1509.4 Ib/hr
Aqueous Phase = 6017.5 Ib/hr |(2746 +
Nonvolatile Organics = 3734.5 Ib/hr |(735 +
Volatile Organics = 3734.5 Ib/hr |(31.4 + 1.
Water = 3734.5 Ib/hr |(1790.1 + 73.7)/2964|
MIBK = 3734.5 Ib/hr |(189.5 + 46.4)/2964| =
Aqueous Phase-3734.5 Ib/hr
Nonvolatile Organics -
1047.0 Ib/hr
Volatile Organics-42.0 Ib/hr
Water - 2348.3 Ib/hr
MIBK - 297.2 Ib/hr
MIBK Phase - 2283 Ib/hr
Nonvolatile Organics -
1005.5 Ib/hr
Volatile Organics-19-7 Ib/hr
Water - 45.6 Ib/hr
MIBK - 1212.2 Ib/hr
218)/4776| = 3734.5 Ib/hr
96)/2964| = 1047.0 Ib/hr
9)/2964| = 42.0 Ib/hr
= 2348.3 Ib/hr
297.2 Ib/hr
148
-------
MIBK Phase = 6017.5 Ib/hr |l812/4776| = 2283.0 Ib/hr
Nonvolatile Organics = 2283.0 Ib/hr |798/1812] = 1005.5 Ib/hr
Volatile Organic = 2283.0 Ib/hr |15.7/1812| =19.7 Ib/hr
Water ='2283.0 Ib/hr |36.2/1812] =45.6 Ib/hr
', •
MIBK =2283.0 Ib/hr 1962.1/18121 = 1212.2 Ib/hr
Evaporator (or Column)
MIBK Soluble Phase -
2283.0 Ib/hr
Nonvolatile Organics -
1005.5 Ib/hr
Volatile Organics -
19.7 Ib/hr
Water - 45.6 Ib/hr
MIBK - 1212.2
Steam
ISO.psia saturated
MIBK = 1212.2 Ib/hr
70°F
358.43°F
hr °'459
BTU
111 Ib • °F
(244
249°F MIBK - 1212.2 Ib/hr
244°F Organics - 1025.2 Ib/hr
244°F Condensate - 355 Ib/hr
- 70) OF
+ 1212.2 Ib/hr |82.5 BTU/lb| + 1212.2 Ib/hr |
0.459 BTU
Ib • °F
(249 - 244) °F | = 199,602 BTU/hr
0.55 BTU
t v
Organics - (estimate cp to be
x ino, 0 ....
) = 1025.2 Ib/hr
0.55 BTU
-, _ OF
•| (244 - 70)°F| = 98,112 BTU/hr
Water = 45.6 Ib/hr [ ]^° .™ I (212 - 70)°F| +45.6 Ib/hr J970.3 BTU/lb |
= 50,721 BTU/hr
Total = 348,435 BTU/hr
Steam Use - x Ib/hr | (1194.1 BTU/lb - 1162.0 BTU/lb) |
+ x Ib/hr | 949. 5 BTU/lb | - 348,435 BTU/hr
x = 355 Ib/hr steam, 150 psia saturated
149
-------
Vacuum Evaporator (Double Effect)—
Water Soluble Phase -
3734.5 Ib/hr
Nonvolatile Organics -
1047.0 Ib/hr
Volatile Organics -
42.0 Ib/hr
Water - 2348.3 Ib/hr
MIBK - 297.2 Ib/hr
Steam
150 psia saturated
Water = 2348.3 Ib/hr
70°F
358.43°F
249°F MIBK - 297.2 Ib/hr
244°F Organics - 1089 Ib/hr
244 °F Condensate - 2817 Ib/hr
(212 - 70)°F| -I-
J.D • r
2348.3 Ib/hr |970.3 BTU/lb| = 2,612,014 BTU/hr
MIBK = 297.2 Ib/hr
+ 297.2 Ib/hr
Ib
(244 - 70)°F|+ 297.2 Ib/hr |82.5 BTlJ/lb1
(249 - 244)°F= 48,937 BTU/hr
Organics - (estimate cp to be °'55 *™ ) = 1089 Ib/hr |0'55 BTU
Ib • °F
•|(244 - 70)°F| = 104,217 BTU/hr
Total = 2,765,169 BTU/hr
Steam Use = x Ib/hr |(1194.1 BTU/lb - 1162.0 BTU/lb)|
+ x Ib/hr|949.5 BTU/lb| = 2,765,169 BTU/hr
»
x = 2,817 Ib/hr steam, 150 psia saturated
Ib
MAJOR EQUIPMENT COST ESTIMATE
Four individual processing schemes have been investigated on the
laboratory scale. Two use raw pyrolytic oil as a feed stock for extraction
while two require that the raw pyrolytic oil undergo a stripping operation
prior to extraction. Two processes employ two stage extraction while two
processes perform a simultaneous extraction in a single stage.
150
-------
The pilot plant was designed so that each of the four processes could
be tested using the single pilot plant installation. For each piece of equip-
ment, the four processes were examined to determine the largest capacity or
size necessary for that particular piece of equipment. For example: process
1-B requires a 1st stage extractor with a volume of 90.96 ft3, while process
2-A requires a volume of 149.2 ft3. Process 2-A was used as the basis for the
design calculations. The pilot plant design basis is a 4 GPM feed rate of raw
pyrolytic oil into the pilot plant system. All pilot equipment is scaled up
directly from experimental results.
Equipment cost estimates are taken from Peters and Timmerhouse [14],
except for estimates of the extractors which are taken from an article by
J. W. Drew [21]. All costs are updated to the period Nov.-Dec. 1979 using the
Chemical Engineering Plant Cost Index. Installations costs are estimated to
be 39% of purchased equipment costs [14]. The evaporators and strippers were
not designed in detail. The heat requirements necessary to perform the par-
ticular ,unit operation were estimated. The results were used directly to
estimate the cost of a piece of equipment that would satisfy the heat require-
ments. The extractor cost estimates are based on Fig. 10, which uses an arbi-
trary column height of 20 feet as a reference point. Although the pilot plant
extractor dimensions would not be expected to be the same as those in the design
calculations, the reference height of 20 feet was used to calculate the equip-
ment cost estimate.
EQUIPMENT COSTS
Pilot Plant—Cost Summary
Raw Oil Storage Tank (1) $ 9,382
Raw Oil Feed Tank (2) 9,382
Vacuum Evaporator (Stripper) (3) 39,090
Extractor (1st Stage) (4) 48,790
Separator (or Holdup Tank) (5) 23,454
Extractor (2nd Stage) (6) 48,790
MIBK Soluble - Holdup Tank (7) 9,382
Evaporator (8) 46,908
MIBK Holdup Tank (9) 3>440
MIBK Soluble - Product Storage Tank (10) 4,691
Water Soluble - Holdup Tank (11) 9,382
Vacuum Evaporator (12) 87,561
Water Soluble - Product Storage Tank (13) 7,193
MIBK Storage Tank (14) 3,440
Volatiles - Product Storage Tank (15) 4,691
Spent Oil - Product Storage Tank (16) 4,691
Water Storage Tank (17) 5,629
Total Installed Equipment Cost 365,896
Instrumentation and controls - (9.35% of
installed equipment cost) o 01,
Piping - (22.3% of installed equipment cost) 81,211
Electrical - (7.2% of installed equipment cost) 26,345
Total Pilot Plant Equipment Cost $508,047
151
-------
Pilot Plant Cost Estimates—Combined Scheme for all Four Continuous
Extraction Processes
Raw Oil Storage Tank—(1)
Use a Tank volume of 500 gal (304ss)
From Figure 13-56 f 14 1 Cost of mixing, storage, and pressure tanks:
Purchased cost = $3000 if^f^f I = $6749
Installed cost = $6749|l.39| = $9382
Raw Oil Feed Tank—(2)
Use a tank volume of 500 gal (304ss)
From Figure 13-56 [ 14] Cost of mixing, storage, and pressure tanks:
Purchased cost = $3000 l^gyl = $6749
Installed cost = $6749 |l.39| = $9,382
Vacuum Evaporator—(Stripper)— (3)
Heat Requirements—From Process 1-B q = 501,464 BTU/hr
From Process 2-B q = 510,979 BTU/hr
Use Process 2-B for design calculations
A^ = (358.43 - 70) °F At, = (220 - 170) °F
A* = Atl ~ At2 = 288.43 - 50
Clm ln(At1/At2) ln(288.43/50) =
q = UAAt, ; estimate U = 200 ^JJL
_ 510,979 _ 2
~ 200(136) - 18'78 ft
From Figure 14-28 [ 14 ] agitated falling-film evaporators (304ss)
Purchased cost = $12,500 if^f^l = $28,122
Installed cost = $28,122 |l.39| = $39,090
Extractor-lst Stage—(4)
Process 1-A
Ra» Oil . 2432.6 Ib/hr ll IL |- 65 »in | - 34.76 ft3
152
-------
Water = 5016.2 Ib/hr \-^ I ^ I 65 min| . 87.09 ft3
Total volume = 121.85 ft3
Process 1-B
Vacu™ Stripped Oil - 2075.0 Ib/hr | - \- _t_ | 65 Mn| . 29.10 ft3
Water - 3563.3 Ib/hr \^^ |^_ | 65 ^ , 61.86 ft3
Total Volume = 90.96 ft3
Process 2-A
Raw Oil - 2470.6 Ib/hr \^^ \^ |^- | 65 min| - 34.76 ft3
3
Water = 3470.3 Ib/hr |- | | 65 min| - 60.25 ft3
MIBK = 2500.4 Ib/hr | ^ |^ |^ | 65 min| = 54.19 ft3
Total Volume = 149.2 ft3
Process 2-B
3
Vacuum Stripped Oil = 2114.2 Ib/hr | *J 1U | -~^ I ,.hr, | 65 min| = 29.65 ft3
O/.4 J.D l.ZJo oO min
Water = 2393.9 Ib/hr | I 65 min = 41'56
MIBK = 1509.4 Ib/hr | - OoT ' 60 65 m±nl = 32'72
Total Volume = 121.85 ft3
Use Process 2-A for design calculations
Use an extractor volume of 150 ft^ (304ss)
From Ref . [21]: V = 7- d h
.where h is assumed to be 20 feet
150 ft3 = ^ d2(20)
d = 3.09 ft or 37.08 inches
From Figure 10 [21] - Cost of Columns:
Purchased cost - $32, 000. | 0.8 | |^|y| | - $35,100
The factor given for converting from a 316ss column to a 304ss column is 0.8.
Installed cost - $35,100 |l.39| - $48,790
153
-------
Separator (or Holdup Tank) — (5)
Use Process 2-A for design calculations
rt
Raw Oil = 2470.6 Ib/hr | ^\ ^ T~234' = 32'085
MIBK - 2500.4 Ib/hr | ^hb ' 0^01 ' = 5°'°26
ft3 , 3
Water = 3470.3 Ib/hr ,. . .. = 55.61 ft /hr
oz . 4 Ib
Total Volume = 137.72 ft3/hr [ 7.48 gal j = 103Q gal/hr
ft
Choose a 3 hour Holdup = 3090 gal
Use a separator volume of 3000 gal (304ss)
From Figure 13-56 [ 14 ] Cost of mixing, storage, and pressure tanks:
Purchased cost = $7500 ~fj = $16,873
Installed cost = $16,873 |l.39J = $23,454
Extractor-2nd Stage-- (6)
Process 1-A
3
Aqueous Phase = 6083.9 Ib/hr | ,.fj .. I . * I ,.hr. I 65 min| = 85.52 ft3
b£. 4 Ib I.ZJD oO mm
3
MIBK = 2436.6 Ib/hr | ** .. |n „., | ,.-hr. | 65 min| = 52.81 ft3
1 62.4 Ib ' 0.801 60 mm ' '
Total volume = 138.34 ft
Process 2-A
3
Aqueous Phase - 4216.7 Ib/hr | lL_ | | _ | 65 min| =59.28 ft3
3
MIBK = 1688.8 Ib/hr ,.fj \ ^ I .-hr. I 65 mini = 36.60 ft3
62.4 Ib ' 0.801 ' 60 mm ' '
Total Volume = 95.88 ft3
3
' Use an extractor volume of 150 ft (304ss)
From Ref . [ 21 J: V = y d2h
4
where h is assumed to be 20 feet
150 ft3 = 2- d2(20)
d = 307 ft2 or 37.08 inches
154
-------
From Figure 10 [ 21 ] - Cost of columns:
Purchased cost = $32,000 | 0.8 1 14|^|| = $35,100
-LU.7 * /
The factor given for converting from a 316ss column to a 304ss column is 0.8.
Installed cost = $35,100 | 1.39J = $48,790
MIBK Soluble—Holdup Tank — ( 7 )
Use a tank volume of 500 gal (304ss)
From Figure 13-56 [ 14 J Cost of mixing, storage, and pressure tanks:
Purchased cost = $3000 |~ 1 " $6749
Installed cost = $6749 |l.39| - $9382
Evaporator (MIBK Phase)-- (8)
Heat Requirements — From Process 1-A q = 423,903 BTU/hr
From Process 1-B q = 292,693 BTU/hr
From Process 2-A q = 816,024 BTU/hr
From Process 2-B q = 348,435 BTU/hr
Use Process 2-A for design calculations
At = (244 - 70) °F At = (358.43 - 249) °F
Atl " At2 174 - 109.43
lm ln(At/At ) ~ ln(174/109.43)
TirnTT
q = UAAt, ; Estimate U = 200 - = -
1m , ,.2 OT,
hr-ft • F
A - q _ 816,024 9Q ,n .2
~ UAt, 200(139.23)
-im
From Figures 14-28 [14] agitated falling-film evaporators (304ss)
Purchased cost = $15,000 if^fl = $33,747
Installed cost = $33,747 |l.39| = $46,908
MIBK Holdup Tank~( 9 )
Use Process 1-A for design calculations
o
MIBK - 2436.6 Ib/hr | I l = 48'75
Capacity - 47.75 ft3/hr J7.48 gal/ft3 | = 364.6 gal/hr
155
-------
Choose a 3 hour Holdup = 1094 gal
Use a tank volume of 1100 gal (C-S)
From Figure 13-56 [ 14 ] Cost of mixing, storage, and pressure tanks:
Purchased cost = $1100 | ' 7 1 = $2475
Installed cost = $2475 |l.39| = $3440
MIBK Soluble - Product Storage Tank — (10)
Use Process 2-B for design calculations
MIBK Soluble Organics = 1025.2 Ib/hr | ^ * lfa Ij-f^jl " 13-3
Capacity =13.3 ft3/hr |7.48 gal/ft | =99.5 gal/hr
Use a tank volume of 150 gal (304ss)
From Figure 13-56 [ 14 ] Cost of mixing, storage, and pressure tanks:
Purchased cost = $1500 ~ = $3375
xuy • /
Installed cost = $3375 |l.39| = $4691
Water soluble-Holdup Tank — (11)
Use a tank volume of 500 gal (304ss)
From Figures 13-56 [ 14 ] Cost of mixing, storage, and pressure tanks:
Purchased cost = $3000 o'? | - $6749
xuy • /
Installed cost = $6749 |l.39| = $9382
Vacuum Evaporator (Water Soluble Phase) — (12)
Heat Requirements — From Process 1-A q = 5,422,062 BTU/hr
Process 1-B q = 3,667,715 BTU/hr
Process 2-A q « 3,892,923 BTU/hr
Process 2-B q = 2,765,169 BTU/hr
Use Process 1-A for design calculations
At^ = (220 - 70) °F At2 = (358.43 - 220) °F
.. Atl " At2 150 - 138.43
lm " ln(At1/At2) ~ In (150/138. 43)
BTTT
q = UA t, ; Estimate U = 500 BTU - =~
lm hr.ft2
q = 5.422.062 _ 2
500(144) °-J rt
156
-------
From Figure 14-28 [ 14 ] Agitated falling-film evaporators (304ss)
Purchased cost = $28,000 if^fl = $62,994
Installed cost = $62,994 |l.39J = $87,561
Water Soluble - Product Storage Tank — (13)
Use Process 2-A for design calculations
Water Soluble Organics = 1698.2 Ib/hr | ^ * | *| = 22.04 ft3/hr
Capacity = 22.04 ft3/hr J7.48 gal/ft3 | = 164.8 gal/hr
Use a tank volume of 300 gal (304ss)
From Figures 13-56 [ 14 ] Cost of mixing, storage, and pressure tanks:
Purchased cost = $2300 ' = $5174
_Luy • /
Installed cost = $5174 |l.39| = $7193
MIBK Storage Tank— (14)
Use a tank volume of 1100 gal (C-S)
From Figure 13-56 [ 14 ] Cost of mixing, storage, and pressure tanks:
Purchased cost - $1100 | ^°'°| = $2475
Installed cost = $2475 |l.39| - $3440
Volatiles-Product Storage Tank—(15)
Use Process 2-B for design calculations
Volatiles—
, ft3 , 1 , 3
Organics = 74.1 Ib/hr [, , ., \ n,.\ = 1.134 ft /hr
f 3 3
Water = 290.3 Ib/hr |fi2" lb| = 4.652 ft /hr
Total Volume - 5.786 ft3/hr |7-48 gal[ =43.28 gal/hr
ft
Choose a 3 hour Holdup = 129.8 gal
Use a tank volume of 150 gal (304ss)
From Figure 13-56 [ 14 ] Cost of mixing, storage, and pressure tanks:
Purchased cost = $1500 y = $3375
Installed cost = $3375 |l.39| = $4691
157
-------
Spent Oil Storage Tank— (16)
Use Process 1-A for design calculations
Insoluble Oil Phase —
Organics = 1026.7 Ib/hr I ^ 4 Ib I 1 234 I " 13>334 ft3/hr
fi-3 ' ' 3
Water = 338.1 Ib/hr A0 . .. | = 5.42 ft /hr
O/.4 ±b .
Total Volume = 18.76 ft3/hr [ 7'48 gal [ = 140 gal/hr
ft
Use a tank volume of 150 gal (304ss)
From Figure 13-56 [ 14 ] Cost of mixing, storage, and pressure tanks:
Purchased cost = $1500 | ?Qg'y| = $3375
Installed cost = $3375 |l.39| = $4691
Water Storage Tank — (17)
Use Process 1-A for design calculations
Water =5016.2 Ib/hr | ,0f* .. I = 80.39 ft3/hr
DZ . 4 Ib
Capacity = 80.39 ft3/hr |7.48 gal/ft3 | = 601.3 gal/hr
Choose a 3 hour Holdup = 1804 gal
Use a tank volume of 200. gal (C-S)
From Figure 13-56 [ 14 1 Cost of mixing, storage, and pressure tanks:
Purchased cost = $1800 - $4050
Installed cost = $4050 1 1.39 1 = $5629
158
-------
APPENDIX C
COMMERCIAL PLANT CALCULATIONS
MAJOR EQUIPMENT—MATERIAL BALANCES
Process 1-A—Raw Oil—Two Stage Extraction
Extractor-lst Stage—
Raw Oil - 9000 Ib/hr
Nonvolatile Organics -
7648.7 Ib/hr
Volatile Organics -
630.6 Ib/hr
Water - 720-7 Ib/hr
Water - 18,558.6 Ib/hr
Aqueous Phase-22,509 Ib/hr
Nonvolatile Organics -
3964.0 Ib/hr
Volatile Organics -
516.7 Ib/hr
Water - 18,028.3 Ib/hr
Insoluble Oil Phase -
5049.6 Ib/hr
Nonvolatile Organics -
3684.7 Ib/hr
Volatile Organics -
114.0 Ib/hr
Water - 1250.9 Ib/hr
Raw Oil
Nonvolatile Organics = 9000 Ib/hr |1698/1998| = 7648.7 Ib/hr
Volatile Organics = 9000 Ib/hr |140/1998] = 630.6 Ib/hr
Water = 9000 Ib/hr |160/1998| = 720.7 Ib/hr
Water = 9000 Ib/hr |4120/1998| = 18,558.6 Ib/hr
Aqueous Phase = 9000 Ib/hr |(4806+191)/1998| = 22,509 Ib/hr
Nonvolatile Organics = 22,509 Ib/hr |(808+72)/4997| = 3964.0 Ib/hr
Volatile Organics = 22,509 Ib/hr |(108.7+6.0/4997| = 516.7 Ib/hr
Water = 22,509 Ib/hr | (3889.3 +113)/49971 = 18,028.3 Ib/hr
159
-------
Insoluble Oil Phase = 9000 Ib/hr 11121/19981 = 5049.6 Ib/hr
Nonvolatile Organics = 5049.6 Ib/hr |818/1121| = 3684.7 Ib/hr
Volatile Organics = 5049.6 Ib/hr j25.3/1121| = 114.0 Ib/hr
Water = 5049.6 Ib/hr |277.7/1121| = 1250.9 Ib/hr
Extractor—2nd Stage
Aqueous Phase-22,509 Ib/hr
Nonvolatile Organics -
3,964.0 Ib/hr
Volatile Organics -
516.7 Ib/hr
Water - 18,028.3 Ib/hr
MIBK - 9014.9 Ib/hr
MIBK Soluble Phase -
9,892.2 Ib/hr
MIBK- 9,014.9 Ib/hr
Organics - 877.3 Ib/hr
Water Soluble Phase -
21,631.7 Ib/hr
Water - 18,028.3 Ib/hr
Organics - 3603.4 Ib/hr
MIBK = 22,509 Ib/hr
1200 ml MIBK i .801g
'2400g Aq Phase ml
= 9014.9 Ib/hr
From laboratory analysis 19.58% of the organics in the Aqueous Phase input
stream were present in the MIBK Soluble Phase, and 80.42% of the Organics
in the Aqueous Phase input stream were present in the Water Soluble Phase.
MIBK Soluble Phase
Organics = 4480.7 Ib/hr |.1958| = 877.3 Ib/hr
Water Soluble Phase
Organics - 4480.7 Ib/hr |.8042| = 3603.4 Ib/hr
Evaporator—
MIBK Soluble Phase -
9892.2 Ib/hr
MIBK - 9014.9 Ib/hr
Organics - 877.3 Ib/hr
70°F
Steam
150 psia saturated
358.43°F
249°F MIBK - 9014.9 Ib/hr
244°F Organics - 877.3 Ib/hr
244°F Condensate - 1598 Ib/hr
160
-------
MIBK = 9014.9 Ib/hr
(244.70)°F | + 9014.9 lb/hr|82.5 BTU/lb[
+ 9014.9 Ib/hr
(249-244)°F| = 1,484,402 BTU/hr
Organlcs - (estimate cp to be
55
-LD
=877.3 Ib/hr
°'55
Xb
•|(244 - 70)°F| = 83.958 BTU/hr
Total = 1,568,360 BTU/hr
Steam Use = x Ib/hr | (1194.1 BTU/lb - 1162.0 BTD/lb) | + x Ib/hr
•|949.5 BTU/lb | = 1,568,360 BTU/hr
x = 1,598 Ib/hr steam 150 psia saturated
Vacuum Evaporator —
Water Soluble Phase -
21,631.7 Ib/hr
Water - 18,028.3 Ib/hr
Organics - 3603.4 Ib/hr
70°F
Steam - 19,943 Ib/hr
150 psia saturated
3-58.43°F
170°F Water - 18,028.3 Ib/hr
220°F Organics - 3603.4 Ib/hr
Water = 18,028.3 Ib/hr | (1134. 2 BTU/lb - 137.97 BTU/lb) |
+ 18,028.3 Ib/hr
J.D
Organics = 3603.4 Ib/hr
(170 - 70)°F| = 19,763,163 BTU/hr
(220 - 70)°F| = 297,281 BTU/hr
Total = 20,060,444 BTU/hr
Steam Use = x Ib/hr | (1194.1 BTU/lb - 1153.4 BTU/lb) | + x Ib/hr
• |965.2 BTU/lb | = 20,060,444 BTU/hr
x = 19,943 Ib/hr steam 150 psia saturated
161
-------
Process 1-B—Vacuum Stripped Oil—Two Stage Extraction
Vacuum Evaporator (Stripper)—
Raw Oil - 9000 Ib/hr
Water - 1053.9 Ib/hr
Organics - 7946.1 Ib/hr
Steam - 1844 Ib/hr
Volatiles - 1323.0 Ib/hr
Water - 1053.9 Ib/hr
Organics - 269.1 Ib/hr
Vacuum Stripped Oil -
7677.0 Ib/hr
Nonvolatile Organics -
7551.2 Ib/hr
Volatile Organics -
125.8 Ib/hr
Raw Oil
Organics = 9000 Ib/hr |.8829| = 7946.1 Ib/hr
Water = 9000 Ib/hr |.117l| = 1053,9 Ib/hr
Volatiles = 9000 Ib/hr |.14?| = 1323.0 Ib/hr
Organics = 1323.0 Ib/hr |.2034| = 269.1 Ib/hr
Water = 1323.0 Ib/hr |.7966| = 1053.9 Ib/hr
Vacuum Stripped Oil = 9000 Ib/hr |.853| = 7677.0 Ib/hr
Nonvolatile Organics = 7677 Ib/hr |1621/1648| = 7551.2 Ib/hr
Volatile Organics = 7677 Ib/hr |27/1648| = 125.8 Ib/hr
Volatiles
11.0 BTU
Water - 1053.9 Ib/hr
Ib • °F
= 1,155,285 BTU/hr
Organics = 269.1 Ib/hr
|(170 - 70)°F|+ 1053.9 Ib/hr|996.2 BTU/hr|
(170 - 70)°F|+ 269.1 Ib/hr J195.5 BTU/lb|
= 66,513 BTU/hr
Vacuum Stripped Oil = 7677.0 Ib/hr
55
Ib • °F
633.353 BTU/hr
(220 - 70)°FJ
Total = 1,855,151 BTU/hr
162
-------
Steam Use =x Ib/hr | (1194.1 BTU/lb - 1153.4 BTU/lb) | + x Ib/hr
•|965.2 BTU/lb| = 1,855,151 BTU/tir
x = 1844 Ib/br steam 150 psia saturated
Extractor-lst Stage—
Vacuum Stripped Oil
7677.0 Ib/hr
Nonvolatile Organics -
7551.2 Ib/hr
Volatile Organics -
125.8 Ib/hr
Water - 13,183.2 Ib/hr
Aqueous Phase-15,600.9 Ib/hr
Nonvolatile Organics -
3344.7 Ib/hr
Volatile Organics-94.1 Ib/hr
Water - 12,162.1 Ib/hr
Insoluble Oil Phase -
5259.3 Ib/hr
Nonvolatile Organics -
4206.5 Ib/hr
Volatile Organics - 31.7 Ib/hr
Water - 1021.1 Ib/hr
Water = 7677.0 Ib/hr |2830/1648| = 13,183.2 Ib/hr
Aqueous Phase = 7677.0 Ib/hr |(3194 + 155)/1648| = 15,600.9 Ib/hr
Nonvolatile Organics = 15,600.9 Ib/hr |(667 + 51)/3349| = 3344.7 Ib/hr
Volatile Organics = 15,600.9 Ib/hr |(19-3 + 0.9)/3349| =94.1 Ib/hr
Water = 15,600.9 Ib/hr ](2507.7 + 103.1)/3349| = 12,162.1 Ib/hr
Insoluble Oil Phase = 7677.0 Ib/hr |1129/1648| = 5259.3 Ib/hr
Nonvolatile Organics = 5259.3 Ib/hr |903/1129| = 4206.5 Ib/hr
Volatile Organics = 5259.3 Ib/hr j6.8/1129) =31.7 Ib/hr
Water = 5259.3 Ib/hr 1219.2/11291 = 1021.1 Ib/hr
163
-------
Extractor-2nd Stage—
Aqueous Phase-15,600.9 Ib/hr
Nonvolatile Organics -
3344.7 Ib/hr
Volatile Organics-94.1 Ib/hr
Water - 12,162.1 Ib/hr
MIBK - 6248.2 Ib/hr
MIBK Soluble Phase -
6813.2 Ib/hr
MIBK - 6248.2 Ib/hr
Organics - 565.0 Ib/hr
Water Soluble Phase -
15,035.9 Ib/hr
Water - 12,162.1 Ib/hr
n-i-oaiTi^Q 9«7T ft 1K/V.1-
MIBK = 15,600.9 Ib/hr
1200 ml MIBK . .801g
2400g Aq Phase ' ml
6248.2 Ib/hr
From laboratory analysis 16.43% of the organics in the Aqueous Phase input
stream were present in the MIBK Soluble Phase, and 83.57% of the organics in
the Aqueous Phase input stream were present in the Water Soluble Phase.
MIBK Soluble Phase
Organics = 3438.8 Ib/hr |.1643| = 565.0 Ib/hr
Water Soluble Phase
Organics = 3438.8 Ib/hr |.835?| - 2873.8 Ib/hr
Evaporator—
MIBK Soluble Phase -
6813.2 Ib/hr
MIBK - 6248.2 Ib/hr
Organics-565.0 Ib/hr
Steam
70°F
358.43°F
0.459
MIBK = 6248.2 Ib/hr \^~^
BTU
Organics - (estimate cp to be
249 °F MIBK - 6248.2 Ib/hr
244°F Organics - 565.0 Ib/hr
244°F Condensate - 1103 Ib/hr
F|+ 6248.2 Ib/hr
= 1,028,836 BTU/hr
565.0 Ib/hr I0'55
Total
•|(244 - 70)°F| = 54,071 BTU/hr
1,082,907 BTU/hr
164
-------
Steam Use = x Ib/hr |(1194.1 BTU/lb - 1162.0 BTU/lb)| + x Ib/hr
•|949.5 BTU/lb| = 1,082,907 BTU/hr
x = 1103 Ib/hr steam 150 psia saturated
Vacuum Evaporator—
Water Soluble Phase -
15,035.9 Ib/hr
Water - 12,162.1 Ib/hr
Organics - 2873.8 Ib/hr
70°F
Steam - 13,490 Ib/hr
150 psia saturated
358.43°F
170°F Water - 12,162.1 Ib/hr
220°F Organics - 2873.8 Ib/hr
Water = 12,162.1 Ib/hr |(1134.2 BTU/lb - ,137.97 BTU/lb)|
+ 12,162.1 Ib/hr |^°.BIp | (170-70)°F| = 13,332,459 BTU/lb
Organics = 2873.8 Ib/hr l-^f-tp | (220 - 70)°F| = 237,089 BTU/hr
Total = 13,569,548 BTU/hr
Steam Use = x Ib/hr |(1194.1 BTU/lb - 1153.4 BTU/lb)| +
-. ' :~J
x Ib/hr |965.2 BTU/lb| = 13,569,548 BTU/hr
x = 13,490 Ib/hr steam, 150 psia saturated
Process 2-A—Raw Oil—Simultaneous MIBK and Water Extraction
Extractor—
Raw Oil - 9000 Ib/hr
Nonvolatile Organics -
7649.9 Ib/hr
Volatile Organics -
630.5 Ib/hr
Water - 719.6 Ib/hr
MIBK - 9108.6 Ib/hr
Water - 12,641.8 Ib/hr
Overhead Effluent -
30,750.4 Ib/hr
Nonvolatile Organics -
7649.9 Ib/hr
Volatile Organics-630.5 Ib/hr
Water - 13,361.4 Ib/hr
MIBK - 9108.6 Ib/hr
Insoluble Oil Phase - 0 Ib/hr
165
-------
Raw Oil
Nonvolatile Organics = 9000 Ib/hr (472/555.3| = 7649.9 Ib/hr
Volatile Organics = 9000 Ib/hr |38.9/555.3| = 630.5 Ib/hr
Water = 9000 Ib/hr |44.4/555.3| = 719.6 Ib/hr
Water = 9000 Ib/hr [780/555.31 = 12,641.8 Ib/hr
MIBK = 9000 Ib/hr J562/55.3| = 9108.6 Ib/hr
Separator—
Overhead Effluent -
30,750.4 Ib/hr
Nonvolatile Organics -
7649.9 Ib/hr
Volatile Organics -
630.5 Ib/hr
Water - 13,361.4 Ib/hr
MIBK - 9,108.6 Ib/hr
Aqueous Phase-18,740.7 Ib/hr
Nonvolatile Organics -
5802.3 Ib/hr
Volatile Organics-384.1 Ib/hr
Water - 12,158.8 Ib/hr
MIBK - 395.5
MIBKFhase-12,009.7 Ib/hr
Nonvolatile Organics -
1847.6 Ib/hr
Volatile Organics -
246.4 Ib/hr
Water - 1202.6 Ib/hr
MIBK - 8713.1 Ib/hr
Aqueous Phase = 30,750.4 Ib/hr [(1153 + 3.3)/1897.3| = 18,740.7 Ib/hr
Nonvolatile Organics = 18,740.7 Ib/hr (358/1156.31 = 5802.3 Ib/hr
Volatile Organics = 18,740.7 Ib/hr [(23.6 + OJ)/1156.3 | = 384.1 Ib/hr
Water = 18,740.7 Ib/hr | (748.3 + 0.9)/1156.31 = 12,158.8 Ib/hr
MIBK = 18,740.7 Ib/hr |(23.1 +0.3)/1156.31 = 395.5 Ib/hr
MIBK Phase = 30,750.4 Ib/hr [741/1897.3| = 12,009.7 Ib/hr
Nonvolatile Organics = 12,009.7 Ib/hr |114/741| = 1847.6 Ib/hr
Volatile Organics = 12,009.7 Ib/hr j15.2/741| = 246.4 Ib/hr
Water = 12,009.7 Ib/hr |74.2/741| = 1,202-6 Ib/hr
MIBK = 12,009.7 Ib/hr |537.6/74l| = 8,713.1 Ib/hr
166
-------
Evaporator (or Column)—
70°F
hr
hr
hr
358.43°F
/^- 0.45
9 BTU
f)/,
249 °F MIBK - 8713.1 Ib/hr
244°F Organics - 2094.0 Ib/hr
244°F Condensate - 3028 Ib/hr
/, 7rA°W 4. S71 ^ 1 IK /tit- .
MIBK Soluble Phase -
12,009.7 Ib/hr
Nonvolatile Organics -
1847.6 Ib/hr
Volatile Organics -
246.4 Ib/hr
Water - 1202.6 Ib/hr
MIBK - 8713.1 Ib/hr
Steam
150 psia saturated
MIBK = 8713.1 Ib/hr
J.D r
182.5 BTU/lb | + 8713.1 Ib/hr ifj^T^T^ 1 (2^9 - 244) °FJ =
1,434,708 BTU/hr
Organics = (estimate cp to be °^ ^ } = 2094'° lb/hr l^5? opU| '
|(244 - 70)°F| = 200,396 BTU/hr
Water = 1202.6 lb/hr |^°.BIp| (212 - 70)°F| + 1202.6 lb/hr | •
[970.3 BTU/lb = 1,337,652 BTU/hr
Total = 2,972,756 BTU/hr
Steam Use = x lb/hr [(1194.1 BTU/lb - 1162.0 BTU/lb)|+
x lb/hr [949.5 BTU/lb| = 2,972,756 BTU/hr
x = 3,028 lb/hr steam, 150 psia saturated
167
-------
Vacuum Evaporator (Double Effect)—
Water Soluble Phase -
18,740.7 Ib/hr
Nonvolatile Organics -
5802.3 Ib/hr
Volatile Organics -
384.1 Ib/hr
Water - 12,158.8 Ib/hr
MIBK - 395.5 Ib/hr
Steam
150 psia saturated
Water = 12,158.8 Ib/hr
70°F
358.43°F
249°F MIBK - 395.5 Ib/hr
244°F Organics - 6186.4 Ib/hr
244°F Condensate
(212 - 70)°F|+ 12,158.8 Ib/hr
• | 970. 3 BTU/lb | = 13,524,233 BTU/hr
MIBK = 395.5 Ib/hr [°90TU| (244 - 70)°F|+ 395.5 Ib/hr
(249 - 244) °F | =
I 82.5 BTU/lb| + 395.5 Ib/hr
65,124 BTU/hr
Organics = (estimate cp to be
10.55 BTUi
6186.4 Ib/hr
iQ.55 BTUi
'Ib • °F I
lib . °F i>
•|(244 - 70)°F| = 592,038 BTU/hr
Total = 14,181,395 BTU/hr
Steam Use = x Ib/hr |(1194.1 BTU/lb - 1162.0 BTU/lb)| +
x Ib/hr 1949.5 BTU/lb | = 14,181,,395 BTU/hr
x = 14,447 Ib/hr steam, 150 psia saturated
Process 2-B—Vacuum Stripped Oil—Simultaneous MIBK and Water Extraction —
Vacuum Evaporator (Stripper)—
168
-------
Raw Oil - 9000 Ib/hr
Organics -7946.1 Ib/hr
Water - 1053.9 Ib/hr
Steam - 1844 Ib/hr
Volatiles - 1323.0 Ib/hr
Organics - 269.1 Ib/hr
Water - 1053.9 Ib/hr
Vacuum Stripped Oil -
7677.0 Ib/hr
Nonvolatile Organics -
7452.8 Ib/hr
Volatile Organics
224.2 Ib/hr
Raw Oil
Organics = 9000 Ib/hr |.8829| = 7946.1 Ib/hr
Water = 9000 Ib/hr |.117l| = 1053.9 Ib/hr
Volatiles = 9000 Ib/hr |.147| = 1323.0 Ib/hr
Organics = 1323.0 Ib/hr |.2034| = 269.1 Ib/hr
Water = 1323.0 Ib/hr |.7966J = 1053.9 Ib/hr
Vacuum Stripped Oil = 9000 Ib/hr |.853| = 7677.0 Ib/hr
Nonvolatile Organics = 7677.0 Ib/hr |1629/1678| = 7452.8 Ib/hr
Volatile Organics = 7677.0 Ib/hr |49/1678| = 224.2 Ib/hr
Volatiles
Water = 1053.9 Ib/hr I*'0 BI^ 1(170 - 70)°F[ + 1053.9 Ib/hr |996.2 BTU/lb|
Organics = 269.1 Ib/hr
= 66,513 BTU/hr
Vacuum Stripped Oil = 7677.0 Ib/hr
Total = 1,855,151 BTU/hr
1,155,285 BTU/hr
°;5j-6I17BTU |(170 - 70)°F| + |269.1 lb/hr| 195.5 BTU/lb|
1(220 - 70)°F| = 633,353 BTU/hr
Steam Use = x Ib/hr | (1194.1 BTU/lb - 1153.4 BTU/lb) |
+ x Ib/hr | 965. 2 BTU/lb | = 1,855,151 BTU/hr
x = 1844 Ib/hr steam, 150 psia saturated
169
-------
Extractor—
Vacuum Stripped Oil -
7677.0 Ib/hr
Nonvolatile Organics -
7452.8 Ib/hr
Volatile Organics -
224.2 Ib/hr
Water - 8692.7 Ib/hr
MIBK - 5481.0 Ib/hr
Overhead Effluent -
21,850.7 Ib/hr
Nonvolatile Organics -
7452.8 Ib/hr
Volatile Organics -
224.2 Ib/hr
Water - 8692.7 Ib/hr
MIBK - 5481.0 Ib/hr
Insoluble Oil Phase - 0 Ib/hr
MIBK = 7677.0 Ib/hr |1198/1678] = 5481.0 Ib/hr
Water = 7677.0 Ib/hr |l900/1678| = 8692.7 Ib/hr
Separator—
Aqueous Phase-13,560.6 Ib/hr
Overhead Effluent -
21,850.7 Ib/hr
Nonvolatile Organics -
7452.8 Ib/hr
Volatile Organics -
224.2 Ib/hr
Water - 8692.7 Ib/hr
MIBK - 5481.0 Ib/hr
Nonvolatile Organics -
3801.8 Ib/hr
Volatile Organics -
152.4 Ib/hr
Water - 8527.1 Ib/hr
MIBK - 1079.3 Ib/hr
MIBK Phase - 8290.1 Ib/hr
Nonvolatile Organics -
3651.0 Ib/hr
Volatile Organics-71.8 Ib/hr
Water - 165.6 Ib/hr
MIBK - 4401.7 Ib/hr
Aqueous Phaee---23^850,7-lb/hr |(2746 + 218)/4776| = 13,560.6 Ib/hr
Nonvolatile Organics = 13,560.6 Ib/hr |(735 + 96)/2964| = 3801.8 Ib/hr
Volatile Organics = 13,560.6 Ib/hr |(31.4 + 1.9)/2964| = 152.4 Ib/hr
Water = 13,560.6 Ib/hr |(1790.1 + 73.7)/2964| = 8527.1 Ib/hr
MIBK = 13,560.6 Ib/hr | (189.5 + 46.4)/2964| = 1079-.3 Ib/hr
170
-------
MIBK Phase = 21,850.7 Ib/hr |1812/4776| = 8290.1 Ib/hr
Nonvolatile Organics = 8290.1 Ib/hr |798/1812] = 3651.0 Ib/hr
Volatile Organics = 8290.1 Ib/hr |15.7/1812| =71.8 Ib/hr
Water = 8290.1 Ib/hr |36.2/1812) = 165.6 Ib/hr
MIBK = 8290.1 Ib/hr |962.1/1812| = 4401.7 Ib/hr
Evaporator (or Column)—
MIBK Soluble Phase -
8290.1 Ib/hr
Nonvolatile Organics -
3651.0 Ib/hr
Volatile Organics -
71.8 Ib/hr
Water - 165.6 Ib/hr
MIBK - 4401.7 Ib/hr
Steam
'150 psia saturated
MIBK = 4401.7 Ib/hr
70°F
r
r
r
358.43°F
0.459
BTU
OLI,
249°F MIBK - 4401.7 Ib/hr
244°F Organics - 3722.8 Ib/hr
244°F Condensate - 1289 Ib/hr
_ 7fA°T7 4- AAD1 7 Th/'hr- •
lib • °F 'v
|82.5 BTU/lb| + 4401.7 Ib/hr
= 724,788 BTU/hr
(249 - 244)°F|
Organics = (estimate cp to be ^ . »F ) = 3722.8 Ib/hr 1]^ . °F I
|(244 - 70)°F| = 356,272 BTU/hr
Water = 165.6 Ib/hr |J-;°.B^| (212 - 70)°F| + 165.6 Ib/hr | •
|(970.3 BTU/lb| = 184,197 BTU/hr
Total = 1,265,257 BTU/hr
Steam Use = x Ib/hr |(1194.1 BTU/lb - 1162.0 BTU/lb)|
+ x Ib/hr |949.5 BTU/lb| = 1,265,257 BTU/hr
x = 1289 Ib/hr steam, 150 psia saturated
171
-------
Vacuum Evaporator (Double Effect)—
Water Soluble Phase -
13,560.6 Ib/hr
Nonvolatile Organics -
3801.8 Ib/hr
Volatile Organics -
152.4 Ib/hr
Water - 8527.1 Ib/hr
MIBK - 1079.3 Ib/hr
Steam
150 psia saturated
Water = 8527.1 Ib/hr
70°F
r
r
r
358.43°F
249°F MIBK - 1079.3 Ib/hr
244°F Organics - 3954.2 Ib/hr
244 °F Condensate - 10,229 Ib/hr
11.0 BTU
Ib • °F
970.3 BTU/lb| = 9,484,693 BTU/hr
(212 - 70)°F| + 8527.1 lb/hr| •
MIBK = 1079.3 Ib/hr
(244 - 70)°F| + 1079.3 Ib/hr |
Ib
|82.5 BTU/lb| + 1079.3 Ib/hr |^'4^90^TU |(249 - 244)°F|
= 177,718 BTU/hr
r, • f ... ^ ,. ^ 0.55 BTU^ ,__. 0 ., ,, 10.55 BTU
Organics = (estimate cp to be — 5^—) = 3954.2 Ib/hr
—~~^ : Ib • F
|(244 - 70)°F| = 378,417 BTU/hr
Total = 10,040,828 BTU/hr
Steam Use = x Ib/hr |(1194.1 BTU/lb - 1162.0 BTU/lb)|
+ x Ib/hr |949.5 BTU/lb| = 10,0401828 BTU/hr
x = 10,229 Ib/hr steam, 150 psia saturated
MAJOR EQUIPMENT COST ESTIMATE
Total installed equipment cost estimates were developed for each of
the four extraction processes. The equipment cost summary for each of the
processes is shown below. Detailed equipment cost estimate calculations
are included for Process 1-B only, as an example.
172
-------
The plant design basis is a 9000 Ib/hr or 14.3 GPM feed rate of raw
pyrolytic oil into the plant. All equipment is scaled up directly from exper-
imental results.
Equipment cost estimates are taken from Peters and Timmerhouse [14]
except for estimates of the extractors which are taken from an article by
J. W. Drew [21]. All costs are updated to the period Nov. - Dec. 1979 using
the Chemical Engineering Plant Cost Index. Installation costs are estimated
to be 39% of purchased equipment costs [14],
The evaporators and strippers were not designed in detail. The heat
requirements necessary to perform the particular unit operation were esti-
mated. The results were used to directly, to estimate the cost of a piece of
equipment that would satisfy the heat requirements. The extractor cost esti-
mates are based on Drew [21], which uses an arbitrary column height of 20 feet
as a reference point. Although the plant extractor dimensions would not be
expected to be the same as those in the design calculations, the reference
height of 20 feet was used to calculate the equipment cost estimate.
EQUIPMENT COSTS
Process IB—(2 Stage Continuous Extraction—Vacuum Stripped Oil)—Cost Summary
Raw Oil Storage Tank 1 $187,631
Raw Oil Feed Tank 2 9,382
Vacuum Evaporator - Raw Oil 3 78,180
Volatiles Storage Tank , 4 73,489
Extractor - 1st Stage 5 77,758
Water Storage Tank 6 46,908
MIBK Storage Tank 7 139,472
Spent Oil Storage Tank 8 7,505
Holdup Tank 9 35,963
Extractor - 2nd Stage 10 80,808
MIBK Soluble - Holdup Tank 11 28,145
Evaporator 12 51,599
MIBK Holdup Tank 13 28,145
MIBK Soluble - Product Storage 14 39,090
Water Soluble - Holdup Tank 15 36,901
Vacuum Evaporator 16 150,105
Water Soluble - Product Storage Tank 17 101,008
Total Installed Equipment Cost $1,172,089
173
-------
Raw Oil Storage Tank — 1
Raw Oil = 9000 Ib/hr I *J .. I T~T| = 118.7 ft3/hr
1 62.4 ID ' 1. 215 '
Capacity = 118.7 ft3/hr ]7.48 gal/ft3] =887.9 gal/hr
Assume a two week supply = 887.9 gal/hr | 24 hr/da | 14 da| = 298,348 gal
Use a tank volume of 300, OOQ gal (304ss)
From Figures 13-59 [14] Storage Tanks:
Purchased cost = $16,000 |3.75 | 246.8/109. 7 1 = $134,986
The factor for converting from C-steel to 304ss is 3.75.
Installed cost = $134,986 I.1.39J = $187,631
Raw Oil Feed Tank — 2
Raw Oil = 9000 Ib/hr [ \ | = 118.7 ft3/hr
Capacity = 118.7 ft3/hr | 7.48 gal/ft3 1' = 887.8 gal/hr
Choose a 4 hour Holdup = 474.8 gal
Use a tank volume of 500 gal (304ss)
From Figures 13-56 [14] Cost of Mixing, storage, and pressure tanks:
Purchased cost - $3000 | 246.8/109.7 | = $6749
Installed cost = $6749 |l.39| = $9382
Vacuum Evaporator (Stripper) — 3
Heat requirements — q = 1,855,151 BTU/hr
At1 = (358.43 - 70) °F At2 = (220 - 170)°F
Atl " At2 288.43-50
1m ln(At1/At2) ln(288.43/50)
q = UAAt. ; Estimate U = 200 ^12
lm hr.ft2.°F,s ,.
A g_ = 1.855,151 = 2
A U t. 200(136) oo.z.rt
lm
From;Figure 14-28 [14] agitated falling-film evaporators (304ss)
Purchas-ed cost = $25,000 | 246.8/109.1\ = $56,244
Installed cost = $56,244 1.39| = $78,180
174
-------
Volatiles - Product Storage Tank— 4
Volatiles—
Organics = 269.1 Ib/hr | ft3/62.4 Ib | 1/1.047| - 4.12 ft3/hr
Water = 1053.9 Ib/hr |ft3/62.4 Ib | = 16.89 ft3/hr
Total Volume = 21.01 ft3/hr J7.48 gal/ft3 | = 157 gal/hr
Assume a 1 week capacity = 157 gal/hr | 24 hr/da | 7 da| = 26,400 gal
Use a tank volume of 26,000 gal (304ss)
From Figures 13-56 [ 14] Cost of mixing, storage, and pressure tanks:
Purchased cost = $23,500 | 246.8/109.7| = $52,870
Installed cost = $52,870 |l.39| - $73,489
Extractor - 1st Stage— 5
Vacuum Stripped Oil = 7677 Ib/hr |ft3/62.4 Ib | 1/1.238 | hr/60 min | 65 min |
= 107.66 ft3
Water'= 13,183.2 Ib/hr |ft3/62.4 Ib | hr/60 min | 65 min| = 228.88 ft3
Use a residence time of 65 min.
3
Total volume = 336.53 ft
3
Use an extractor volume of 350 ft (304ss)
IT 2
From Reference [ 21] : V = 7- d h
where h is assumed to be 20 feet
350 ft3 = (Tr/4)d2(20)
d = 4.72 ft or 56.64 inches
From Figure 10 [ 21] - Cost of columns:
Purchased cost = $51,000 | 0.8 | 246.8/1801 = $55,941
The factor given for converting from a 316ss column to a 304ss column is 0.8.
Installed cost = $55,941 |l.39| = $77,758
175
-------
Water Storage Tank— 6
Water = 13,183.2 Ib/hr |ft3/62.4 Ib| = 211.27 ft3/hr
Capacity = 211.27 ft3/hr | 7.48 gal/ft3| =1580 gal/hr
Assume a 3 day supply = 1580 gal/hr | 24 hr/da | 3 da| = 113,781 gal
Use a tank volume of 110,000 gal (C-S)
From Figures 13-56 [14] Cost of mixing, storage, and pressure tanks:
Purchased cost = $15,000 |246.8/109.7| = $33,747
Installed cost = $33,747 |l.39| = $46,908
MIBK Storage tank— 7
MIBK = 6248.2 Ib/hr | ft3/62.4 Ib |1/0.8011 = 125.01 ft3/hr
Capacity = 125.01 ft3/hr |7.48 gal/ft3| = 935 gal/hr
Assume a 1% loss of MIBK in system through pumps, leakage, etc., which
requires makeup.
Assume a two week supply = 935 gal/hr | 24 hr/da | 14 da | .Ol| = 3142 gal
Use a tank volume of 3500 gal (C-S)
From Figures 13-56 [ 14 ] Cost of mixing, storagej and pressure tanks:
Purchased cost = $2400 | 246.8/109.71 = $5400
Installed cost = $5400 |l.39| = $7505
Spent Oil Storage Tank— 8
Insoluble Oil Phase
Organics = 4238.2 Ib/hr |ft3/62.4 Ib |1/1.235| = 55 ft3/hr
Water = 1021.1 Ib/hr |ft3/62.4 Ib| = 16.36 ft3/hr
Total Capacity = 71.36 ft3/hr J7.48 gal/ft3|= 533.77 gal/hr
Assume a one week capacity = 533.77 gal/hr 24 hr/da | 7/da = 89,673 gal
Use a tank volume of 90,000 gal (304ss)
From Figures 13-56 [ 14 ] Cost of mixing, storage, and pressure tanks:
Purchased cost = $44,600 |246.8/109.7| = $100,340
Installed cost = $100,340 |l.39| = 139,472
176
-------
Holdup Tank— 9
Aqueous Phase = 15,600.9 Ib/hr | ft3/62.4 Ib |1/1.235] - 202.44 ft3/hr
Capacity = 202.44 ft3/hr |7.48 gal/ft3| = 1514.3 gal/hr
Choose a 4 hour Holdup = 6057 gal
Use a tank volume of 600 gal (308ss)
From Figures 13-56 [ 14 ] (Cost of mixing, storage, and pressure tanks:
Purchased cost = $11,500 | 246.8/109.7| = $25,872
Installed cost = $25,872 |l.39| = $35,963
Extractor - 2nd Stage— 10
Aqueous Phase = 15,600.9 Ib/hr |ft3/62.4 Ib |1/1.235 |hr/60 min | 65 minj
= 219.31 ft
MIBK = 6248.2 Ib/hr |ft3/62.4 Ib |1/0.801 | hr/60 min | 65 min| = 135.43 ft3
Use a residence time of 65 min
3
Total volume = 354.74 ft
3
Use an extractor volume of 375 ft (304ss)
2
From Reference [ 21 J : V = (ir/4)d h
where h is assumed to be 20 feet
375 ft3 = (TT/4)d2(20)
d = 4.89 ft or 58.63 inches
From Figure 10 [ 21J - Cost of columns:
Purchased cost = $53,000 | 0.8 | 246.8/1801 = $58,135
The factor given for converting from a 316ss column to a 304ss column is 0.8.
Installed cost - $58,135 |l.39| = $80,808
MIBK Soluble - Holdup Tank— 11
A
MIBK = 6248.2 Ib/hr |ft3/62.4 Ib | l/0.80l| = 125.01 ft /hr
Organics = 565 Ib/hr | ft3/62.4 Ib | 1/1.235] =7.33 ft3/hr
177
-------
Capacity = 132.34 ft3/hr | 7.48 gal/ft3 | = 990 gal/hr
Choose a 4 hour Holdup = 3960 gal
Use a tank volume of 4000 gal (304ss)
From Figure 13- 56 [14] Cost of mixing, storage, and pressure tanks:
Purchased cost = $9,000 | 246. 8/109. 7 | = $20,248
Installed cost = $20,248 |l.39| = $28,145
Evaporator (MIBK Phase) — 12
Heat Requirements— q = 1,082,907 BTU/hr
A^ = (244 - 70) °F At2 = (358.43 - 249) °F
At. - At_ 1 74 -
AI-
fl
lm ln(At1/At2) ln(174/109.43)
q = UAAt, ; Estimate U = 200 - -
lm hr • ft2 • °F
A - a - 1*082.907 _ 2
A UAt. " 200(139) Ja-y:) "
lm
From Figures 14-28 [14] agitated falling film evaporators:
Purchased cost = $16,500 1 246. 8/109. 7 | = $37,121
Installed cost = $37,121 |l.39| = $51,599
MIBK Holdup Tank— 13
MIBK = 6248.2 Ib/hr |ft3/62.4 Ib | 1/0. 801 1 = 125.01 ft3/hr
Capacity = 125.01 ft3/hr |7.48 gal/ft3) = 935.1 gal/hr
Choose a 4 hour Holdup = 3740 gal
Use a tank volume of 4000 gal (304ss)
From Figure 13-56 [14] Cost of mixing, storage, and pressure tanks:
Purchased cost = $9,000 | 246.8/109.7 1 = $20,248
Installed cost = $20,248 |l.39| = $28,145
MIBK Soluble - Product Storage Tank — 14
MIBK Soluble Organics = 565 Ib/hr ft3/62.4 Ib | 1/1.235] = 7.33 ft3/hr
Capacity =7.33 ft3/hr | 7.48 gal/f t3 | =54.8 gal/hr
178
-------
Assume 1 week capacity = 54.8 gal/hr (24 hr/da | 7 da/wk| = 9211 gal
Use a tank volume of 9,000 gal (304ss)
From Figure 13-56 [14] Cost of mixing, storage, and pressure tanks:
Purchased cost = $12,500 | 246.8/109.7] = $28,122
Installed cost = $28,122 | 1.39| = $39,090
Water Soluble - Holdup Tank— 15
Water = 12,162.1 Ib/hr |ft3/62.4 lb| = 194.9 ft3/hr
Organics = 2873.8 Ib/hr | ft3/62.4 lb| 1/1.235) = 37.29 ft3/hr
Capacity = 232.2 ft3/hr | 7.48 gal/ft3] = 1736.8 gal/hr
Choose a 4 hour Holdup = 6947 gal
Use a tank volume of 7,000 gal (304ss)
From Figure 13-56 [ 14] Cost of mixing, storage, and pressure tanks:
Purchased cost = $11,800 | 246.8/109.7 | = $26,547
Installed cost = $26,547 |l.39| = $36,901
Vacuum Evaporator — 16
Heat Requirements — q = 13,569,548 BTU/hr
At-L = (220 - 70)°F At2 = (358.43 - 220)°F
Atl"At2 150- 138.43 _ ,.„
Atlm ln(At1/At2) " ln(150/138.43)
q = UAAtn ; Estimate U = 500 - = -
lm hr-ft -°F
A = q/UATlm = 13,569,548/500(144) = 188.46 ft2
From Figure 14-28 [14] agitated falling-film evaporators
Purchased cost = $48,000 | 246.8/109.7 | = $107,989
Installed cost = $107,989 |l.39| = $150,105
Water Soluble - Product Storage Tank — 17
q O
Water Soluble Organics = 2873.8 Ib/hr | ft /62.4 Ib | 1/1.235] = 37.29 ft ,/hr
179
-------
Capacity = 37.29 ft3/hr |7.48 gal/ft3| = 279 gal/hr
Assumg a 1 week capacity = 279 gal/hr | 24 hr/da | 7 da/wk| = 46,860 gal
Use a tank volume of 50,000 gal (304ss)
From Figure 13-56 [14] Cost of mixing, storage, and pressure tanks:
Purchased cost • $32,300 | 246.8/109.7] = $72,668
Installed cost = $72,668 I 1.39| - $101,008
180
-------
APPENDIX D
PHYSICAL PROPERTIES
TABLE D-l. TYPICAL VOLATILES ANALYSIS*
Component
Water
Acetic Acid
Methanol
Furfural
Formic Acid
Propionic Acid
Unknown
Weight
Per cent
68.24
20.48
1.70
2.00
2.42
0.60
4.56
*
Experimental Results
TABLE D-2. HEAT CAPACITY ESTIMATION* - VOLATILES
Component
Water
Acetic Acid
Methanol
Furfural
Formic Acid
Propionic Acid
Unknown
cp
(A)
Weight
Per cent
68.24
20.48
1.70
2.00
2.42
0.60
4.56
- Weighted Average -
(B)
cp
BTU
Ib • °F
1.00
0.522
0.590
0.416
0.524
0.560
0.50 (Est.)
(A) * (B)
0.6824
0.1069
0.0100
0.0083
0.0127
0.0034
0.0228
0.8465
[ 15 ] Table 3-176 Specific Heats of Organic Liquids
181
-------
TABLE D-3. HEAT CAPACITY ESTIMATION—-VOLATILES LESS WATER
Component
Methanol
Formic Acid
Acetic Acid
Propionic Acid
Furfural
Unknown
Total
grams
1.4
2.0
16.9
0.5
1.65
3.75
26.2
(C)
cp
BTU
Ib • °F
0.590
0.524
0.522
0.560
0.416
0.50 (Est.)
(D)
Weight
Fraction
0.0534
0.0763
0.6450
0.0191
0.0630
0.1431
(C) * (D)
0.0315
0.0400
0.3367
0.0107
0.0262
0.0716
cp - weighted average = 0.5167
Experimental Results
TABLE D-4. DENSITY ESTIMATION—VOLATILES LESS WATER
Component
Acetic -Acid
Methanol
Furfural
Formic Acid
Propionic Acid
Unknown
(A)
Weight
Fraction
0.6450
0.0534
0.0630
0.0763
0.0191
0.1431
(B)
Density
g/ml
1.0491
0.7914
1.1598
1.220
0.992
1.00 (Est.)
•-.v'-:."
(A) * (B)-"';-1*'
-/i...-
0.6767
0.0423
0.0731
0.0931
0.0189
0.1431
Density (p) - Weighted Average = 1.0472
TABLE D-5. HEAT OF VAPORIZATION* ESTIMATION—VOLATILES LESS WATER
Component
Acetic Acid
Methanol
Furfural
Formic Acid
Propionic Acid
Unknown
Boiling
Point- 0;C
118.3
64.7
60.5
101
139.5
—
Heat of Vaporization
cal/g BTU/lb
96.8
262.8
107.5
120
98.8
100 (Est)
174.24
473.04
193.5
216.0
177.84
180 (Est)
Weight
Per cent
0.6450
0.0534
0.0630
0.0763
0.0191
0.1431
112.3848
25.2603
12.1905
16.4808
3.3967
25.7580
Heat of Vaporization - Weighted Average = 195.5 BTU/lb
[20] , pp. 9-85 — 9.95
182
-------
TABLE D-6. SPECIFIC HEATS OF VARIOUS ORGANIC COMPOUNDS*
Compound
Methyl Ethyl Ketone
Methyl Isobutyl Ketone
Methyl Isopropyl Ketone
Methyl Hexyl Ketone
Water
Specific Heat
BTU CAL
Ib • °F "1 g • °F
0.549
0.459
0.525
0.55
1.0
Temperature
Range (°C)
20 - 78°C
20
20 - 91
22 - 168
15
[ 20 ],pp. 9-133
TABLE D-7. HEAT OF VAPORIZATION OF VARIOUS ORGANIC COMPOUNDS
Compound
Methyl Ethyl Ketone
Methyl Isopropyl Ketone
Methyl n-Butyl Ketone
Methyl n-Amyl Ketone
Methyl Hexyl Ketone
Water
Heat of Vaporization
cal/gram
106.0
89.8
82.4
82.7
74.1
539.55
Temperature
°C
78.2
92
127
149.2
173
100
[ 20 ], pp. 9-91
TABLE D-8. PROPERTIES OF MIBK*
Molecular Weight
Density g/ml
Melting Point
Boiling Range
Solubility in Water
100.16
0.801 @20°/408
-84.7 °C
117 - 119 °C
2g/100g Water at 20°C
[20], pp. 7-54
183
-------
GLOSSARY
dissolved organics: The nonvolatile material remaining after evaporation of
solvent of a fraction or phase.
fraction: A solution or solid derived from extracting a phase (as defined
below) with an immiscible solvent.
neutrals of high aromaticity (NHA): Compounds in the pyrolytic oils which
are nonpolar and nonacidic and exhibit UV fluorescing and absorbing
(254 nm) characteristics.
nonvolatile organics (NVO): The fraction of organic material remaining after
vacuum stripping at approximately 2 mm Hg and ambient temperature
which contains phenolics, polyhydroxy neutral compounds, and neu-
tral compounds of high aromaticity. «&*.
organic volatiles: The organic volatiles is equal to the difference between
the total volatiles and amount of water in the total volatiles. !'!
•"•j
phase: A solution or solid derived from the original pyrolytic oil sample
by evaporation or extraction, e.g. volatile phase, aqueous phase,
organic phase, insoluble tar.
- - i
phenolics: The class of acidic compounds which are titratable with meth- '
anolic potassium hydroxide in N,N-dimethyl formamide solvent, and
identifications are confirmed by GC, LC, TLC and IR evidence.
polyhydroxy compounds: The class of nonacidic compounds which are very
water soluble, produce a blue color with Orcinol reagent, which is
a characteristic of sugars, and have RF values similar to those of
known sugar on a TLC plate.
total volatiles: The total volatile material, including both water and
organics, removed by vacuum stripping at approximately 2 mm Hg and
ambient temperature.
organic volatiles: The organic volatiles is equal to the difference between
the total volatiles and amount of water in the total volatiles.
184
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TECHNICAL REPORT DATA
{Please read Instructions on the reverse before completing)
1. REPORT NO.
EPA-600/2-80-122
3. RECIPIENT'S ACCESSION NO.
4. TITLE AND SUBTITLE
5. REPORT DATE
Pyrolytic Oils - Characterization and Data Development
for Continuous Processing
August 1980 (Issuing Date)
6. PERFORMING ORGANIZATION CODE
. AUTHOR(S)
]. A. Knight, L. W. Elston, D. R. Hurst, and
R. J. Kovac
8. PERFORMING ORGANIZATION REPORT NO.
9. PERFORMING ORGANIZATION NAME AND ADDRESS
Engineering Experiment Station
Georgia Institute of Technology
Atlanta, Georgia 30332
10. PROGRAM ELEMENT NO.
IDr.Ria
11. CONTRACT/GRANT NO.
R804416 and R806403
12. SPONSORING AGENCY NAME AND ADDRESS
Municipal Environmental Research Laboratory— Cin., OH
Office of Research and Development
U.S. Environmental Research Agency
Cincinnati, Ohio 45268
13. TYPE OF REPORT AND PERIOD COVERED
14.'SPONSORING AGENCY CODE
EPA/600/14
15. SUPPLEMENTARY NOTES
Project Officer: Charles J. Rogers (513) 684-4335
16. ABSTRACT
Pyrolytic oils produced by the pyrolysis of forestry residues in a vertical bed,
countercurrent flow reactor have been thoroughly characterized. The pyrolytic oils
were produced in a 500-1b. per hour pilot plant and in a 50-ton per day field
development facility. The overall chemical and physical properties have been de-
termined by standard analytical techniques. The oils are dark brown to black with
a burnt, pungent odor and have a boiling range of about 100°C to approximately 200°C
at which point thermal degradation begins to occur. Pyrolytic oils contained
phenolics, polyhydroxy neutral compounds and volatile acidic compounds.
7.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.IDENTIFIERS/OPEN ENDED TERMS
COSATI Field/Group
Pyrolytic oils
Pyrolysis
Polyhydroxy neutral compounds
Degradation
Extraction
Phenols
thermal degradation
volatilization and
compounds
13B
8. DISTRIBUTION STATEMENT
Release to public
19. SECURITY CLASS (This Report)
Unclassified
11. NO. OF PAGES
197
20. SECURITY CLASS (This page)
Unclassified
22. PRICE
EPA Form 2220-1 (9-73)
185
U.S. GOVERNMENT PRINTING OFFICE: 1980—657-165/0053
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