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dissertation entitled
Metabolic Engineering of Product Formation During Carbon

Monoxide Fermentation by Butyribacterium methylotrophicum

 

presented by

Andrew Jacob Grethlein

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METABOLIC ENGINEERING OF PRODUCT FORMATION DURING CARBON
MONOXIDE FERNIENTATION BY B UTYRIBACT ER] UM MET HYLOTROPHIC UM

By

Andrew Jacob Grethlein

A DISSERTATION

Submitted to
Michigan State University
in partial fulﬁllment of the requirements
for the degree of

DOCTOR OF PHILOSOPHY

Department of Chemical Engineering
199 l

ABSTRACT
METABOLIC ENGINEERING OF PRODUCT FORMATION DURING CARBON
MONOXIDE FERMENTATION BY BUTYRIBACT ERI UM MET HYLOTROPHICUM
By
Andrew Jacob Grethlein

The regulation of carbon monoxide (CO) gas metabolism was examined in the
anaerobic bacterium Butyribacten'um methylotrophicum using continuous fermentation and
batch bottle studies. A cell recycle system was integrated with a chemostat to increase cell
density and bias the cell population towards the stationary (resting) state. Culture media
was optimized to support high cell densities and maintain CO as a limiting substrate. This
system was operated at fermentation pH values between 5.1 and 7.2. Results indicated
three metabolic regimes for this system. Between pH values of 6.0 and 7.2, steady-state
operation was achieved, and a 5-10 fold increase in acid and alcohol production was
observed relative to chemostat operation. Between pH values of 5.5 and 6.0, oscillations
in butyrate and acetate production were observed, with butanol as the major product over a
72 hour period during an initial start-up period. Between pH values of 5.1 and 5.5, the
cells exhibited a death phase response.

Under conditions of decreasing pH and high acetate concentration in batch culture, B.
methylotrophicum exhibited consumption of both CO and acetate, or CO and butyrate,
with net production of butyrate or acetate, butanol and ethanol. The tolerance to both
ethanol and butanol was less than 5 g/L. Hydrogen sulﬁde was toxic above partial
pressures of 2%. These studies indicate that regulation of CO metabolism in B.
methylotrophicwn can be achieved using a variety of environmental conditions.

A metabolic model, based on a linear system of equations for the known pathways of
CO metabolism, was developed to analyze the fermentation data. The model was accurate

compared to other methods of data analysis. The metabolic analysis showed decreasing

ATP yields with decreasing fermentation pH for steady-state cultures. For unsteady-state
results, oscillations in ATP production were shown concommitant with oscillations in
substrate consumption and product formation. From the modeling results, low
fermentation pH was concluded to cause a decrease in the net ATP yields between steady-
state and unsteady state systems, primarily via decreases in acetate production, inhibition

of cell formation, and by decreasing the overall rate of CO metabolism.

ACKNOWLEDGMENTS

This work would not have been possible without the unwavering support of the staff at
the Michigan Biotechnology Institute, and I wish to express my deepest appreciation to all
of the employees and associates at MBI. I am particularly indebted to both Dr. Rathin Datta
and Dr. Mahendra Jain for their advice and support during the course of this work. Special
thanks also goes out to all of my friends and peers at both MBI and MSU, including Detlef
Schartges, Ali Siahpush, and Cathy McGowan. I would especially like to thank Greg Reid
for his steadfast friendship during the last 5 years, as well as for frequent reality checks.
Finally, the success of this entire undertaking was due in large part to my thesis advisor,
Dr. R. Mark Worden, who allowed me the freedom to pursue my own interests, and who

consistently supported my efforts with advice, critique, and most of all, with enthusiasm.

iv

TABLE OF CONTENTS

LIST OF TABLES: ................................................................... vii
LIST OF FIGURES: .................................................................... ix
INTRODUCTION: ..................................................................... 1
CHAPTER 1: TECHNICAL AND SCIENTIFIC REVIEW: ........................ 5
1.1. Synthesis Gas Chemical Technology: .......................................... 5
1.1.1. The Case for Coal Utilization: .......................................... 5
1.1.2. Direct versus Indirect Liquefaction: ................................. 7
1.1.3. Gasiﬁcation and Synthesis Gas Production: ........................ 9
1.1.4. Puriﬁcation of Synthesis Gas: ........................................ 14
1.1.5. Synthesis Gas Utilization: ........................................ 22
1.1.6. Application of Synthesis Gas-Derived Compounds: ............. 34
1.1.7. Future Prospects: ................................................. 37
1.2. Microbial Metabolism of Synthesis Gas Components: ......................... 38
1.2.1. Principles of l-Carbon Metabolism: ............................... 3 8

1.2.2. Carbon Monoxide, Carbon Dioxide, and Hydrogen
Metabolizing Bacteria: ................................................. 41
1.2.3. Principles of Acetogenic Metabolism: ............................... 47
1.2.4. Potential Application of Acetogenic Bacteria: ...................... 5 3
1.3. Physiology and Metabolism of Butyn’bacterium methylotrophr'cum: . . . .5 6
1.3.1. Origin and General Characteristics of the CO Strain: ............. 56
1.3.2. Heterotrophic Metabolic Characteristics: ...................... 57
1.3.3. Autotrophic Metabolic Characteristics: ...................... 59
1.4. Biological Processing of Coal-Derived Gases: ............................... 63
1.4.1. Potential Advantages of Biological Processing: ............. 64
1.4.2. Potential Disadvantages of Biological Processing: ............. 65
1.4.3. The MBI Indirect Liquefaction Process: ...................... 66
1.4.4. Objectives and Scope of the Present Work: ...................... 70
CHAPTER 2: EXPERIMENTAL MATERIALS AND METHODS: ............. 7 2
2.1. Microorganism and Culture Conditions: ........................................ 72
2.2. Culture Media and Gases: ................................................. 72
2.3. Fermentation Equipment: ................................................. 74
2.4. Fermentation Conditions: ................................................. 7 6

2.5. Fermentation Broth Analysis: ................................................. 77

2.6. Fermentation Product Analysis: ................................................. 7 8

CHAPTER 3: METABOLIC TOLERANCE AND REGULATION STUDIES: .80

3.1. Objectives: ................................................................... 80

3.2. Effect of Alcohols on Growth and Product Formation: ...................... 81

3.3. Effect of Acids on Growth and Product Formation: ...................... 90

3.4. Effect of Sulﬁde on Growth and Product Formation: ..................... 100

3.5. Media Development Studies: ................................................ 1 10

CHAPTER 4: CELL RECYCLE FERMENTATION STUDIES: ..................... 114

4.1. Objectives: .................................................................. 1 14

4.2 System Set-up and Design: ................................................ 1 17

4.3. Results for Continuous Cell Recycle CO Fermentations: ............ 120

4.3.1. Steady-State Results: ................................................ 121

4.3.2. Unsteady-State Results: ....................................... 130

4.3.3. Death Phase Results: ................................................ 137

4.3.4. Cosubstrate Results: ................................................ 137

4.4. Summary Discussion of Fermentation Results: .............................. 13 8

CHAPTER 5: METABOLIC MODELING AND

FERMENTATION ANALYSIS: ....................................... 140

5.1. Objectives and Principles of the Model: ....................................... 140

5.2. Development of the Metabolic Model: ....................................... 141

5.3. Application to Chemostat Results: ....................................... 148

5.4. Application to Steady-State Cell Recycle Fermentation Results: ............ 15 3
5.4.1. Case 1: Uniquely Speciﬁed System;

No Cell Mass Production: .............................. 15 3

5.4.2. Case 2: Uniquely Speciﬁed System: .............................. 160

5.4.3. Cases 3-5: Overspeciﬁed Systems: .............................. 166

5.5. Application to Unsteady-State Cell Recycle Fermentation Results: . ..168

5.6. Limitations of the Model: ................................................ 177

CHAPTER 6: CONCLUSIONS AND RECOMMENDATIONS: ............ 180

APPENDIX: ........................................................................... 1 84

LIST OF REFERENCES: ......................................................... 1 87

vi

LIST OF TABLES

Table 1. Synthesis gas compositions from commercial and near-commercial coal

gasiﬁcation processes: ....................................................... 13
Table 2. CO/Hz ratios of direct synthesis routes for

chemicals from synthesis gas: .............................................. 30
Table 3. Properties of various fuel oxygenates: ..................................... 3 6

Table 4. Reactions in the unicarbonotrophic and autotrophic
pathways for acetyl-CoA synthesis in the acetogens

B. methylotrophicum and C. thermoaceticum: ............................ 50
Table 5. l-Carbon catabolic transformations observed

with B. methylotrophicum: ....................................................... 60
Table 6. Components of the phosphate buffered basal medium: ................... 7 3
Table 7. Balanced carbon and available electron equations for batch incubation in

the presence of ethanol during growth on C0 gas: ............................ 84

Table 8. Balanced carbon and available electron equations for batch incubation in
the presence of butanol during growth on C0 gas: ............................ 84

Table 9. Effect of increasing alcohol concentration on cell density and product
formation during growth on C0 gas: .............................................. 8 5

Table 10. Effect of increasing alcohol concentration on cell growth and CO
consumption during growth on C0 gas: ..................................... 85

Table 11. Balanced carbon and available electron equations
for batch incubation of stationary phase cells in the

presence of alcohols during growth on C0 gas: ............................ 87
Table 12. Balanced carbon and available electron equations for batch incubation in

thepresence of acetate during growth on C0 gas: ............................ 95
Table 13. Balanced carbon and available electron equations for batch incubation in

the presence of butyrate during growth on C0 gas: ............................ 95

Table 14. Theoretical molar ATP and electron yields from substrate-level
phosphorylation for products derived during CO metabolism in
Butyribacteriwn methylotrophicwn: .............................................. 98

Table 15. Experimental pH proﬁles showing dynamics of H28 dissolution: ......... 102

vii

Table 16. Effect of liquid phase sulﬁde concentration on cell growth of B.
methylotrophicum grown on C0 gas at an initial pH of 7.0: .................. 109

Table 17. Continuous, steady-state CO fermentation stoichiometries as a function
of pH for B. methylotrophicum: ............................................. 1 16

Table 18. Continuous, steady- state CO fermentation product concentrations as a
function of pH for B. methylotrophicum: .................................... 1 16

Table 19. Overall culture response as a function of fermentation pH during
cell recycle fermentation: ...................................................... 121

Table 20. Steady-state fermentation stoichiometries as a function of pH for
continuous fermentation of CO gas: ............................................. 122

Table 21. Continuous, steady-state CO fermentation molar yield coefficents as a
function of pH for B. methylotrophicum: .................................... 122

Table 22. Continuous, steady-state CO fermentation carbon partitioning as a
function of pH for B. methylorrophicum: .................................... 124

Table 23. Continuous, steady-state CO fermentation electron partitioning as a
function of pH for B. methylotrophicum: .................................... 124

Table 24. Continuous, steady-state CO fermentation parameters, product
concentrations, and CO consumption rates as a function of pH
for B. methylotrophicum: ...................................................... 1 26

Table 25. Continuous, steady-state CO fermentation CO consumption
rates as a function of fermenter volume and culture

OD for B. merhylotrophicum: ............................................. 1 26
Table 26. Unsteady-state, initial fermentation stoichiometries as a function of pH

for cell recycle fermentation of CO gas: .................................... 13 3
Table 27. Continuous, initial unsteady-state CO fermentation carbon partitioning

as a function of pH for B. methylotrophicum: ........................... 136
Table 28. Continuous, initial unsteady-state CO fermentation electron partitioning

as a function of pH for B. methylotrophicum: ........................... 136
Table 29. Continuous, steady-state CO fermentation carbon and electron partitioning

at a pH of 5.9 in the presence of acetate for B. methylotrophicum: ......... l3 8
Table 30. Standard free energies of reaction for CO conversion: .................. 139

Table 31. Comparison of CO:COz ratios from chemostat culture between metabolic
model and carbon and electron balancing calculations: .................. 151

Table 32. Comparison of CO:COz ratios from steady-state cell recycle culture

between metabolic model and carbon and electron balancing calculations
for Case 2: ........................................................................ 161

viii

LIST OF FIGURES

Figure 1. Schematic of typical sulfur recovery and removal operations during

catalytic conversion of coal—derived synthesis gas: ............................ 21
Figure 2. Potential industrial uses of fuels and chemicals from coal-derived

synthesis gas: ................................................................ 23
Figure 3. Speciﬁc uses for l-carbon compounds from coal-derived

synthesis gas: ................................................................ 27
Figure 4. General process scheme for SASOL Fischer-Tropsch synthesis plant,

Sasolburg, South Africa (Dry, 1987): ..................................... 29
Figure 5. General pathway for unicarbonotrophic, acetogenic metabolism: .......... 49

Figure 6. General mechanism for electron transport-coupled
phosphorylation (ETP): ....................................................... 52

Figure 7. Potential biological routes for fuels and chemicals production
from synthesis gas components: .............................................. 54

Figure 8. Proposed bioconversion scheme for production of fuel alcohols
from coal-derived synthesis gas: .............................................. 68

Figure 9. Effect of increasing ethanol concentration during growth of
B. methylotrophicum on C0 gas: .............................................. 82

Figure 10. Effect of increasing butanol concentration during growth of
B. methylotrophicum on C0 gas: .............................................. 82

Figure 11. Effect of increasing ethanol concentration during growth of
B. methylotrophicum on C0 gas: .............................................. 8 6

Figure 12. Effect of increasing butanol concentration during growth of
B. methylorrophicum on C0 gas: .............................................. 86

Figure 13. Effect of increasing acetate concentration during growth of
B. methylotrophicwn on C0 gas: .............................................. 9 3

Figure 14. Effect of increasing butyrate concentration during growth of
B. methylotrophicum on C0 gas: .............................................. 94

Figure 15. Effect of increasing NaCl concentration during growth of
B. methylotrophicum on C0 gas: .............................................. 97

ix

Figure 16. Effect of initial H28 gas phase concentration dming growth of
B. methylotrophicum on C0 gas: ............................................. 104

Figure 17. Effect of increasing H28 concentration on acetate production
during growth of B. methylotrophicum: .................................... 104

Figure 18. Acetate yields as a function of initial H28 concentration
during growth of B. merhylotrophicum on C0 gas: ........................... 105

Figure 19. Effect of initial sulﬁde concentration during growth of
B. methylotrophicum on C0 gas: ............................................. 107

Figure 20. Effect of initial sulﬁde concentration during growth of
B. methylotrophicum on C0 gas: ............................................. 107

Figme 21. Speciﬁc and molar yields as a function of initial sulﬁde concentration
during growth of B. methylotrophicum on C0 gas: ........................... 108

Figure 22. Speciﬁc and molar yields for product formation during
growth of B. methylotrophicum on C0 gas: .................................... 108

Figure 23. Effect of media supplements on the growth
of B. methylotrophicum on C0 gas: ............................................. 1 13

Figure 24. Effect of increasing yeast extract concentration on the growth
of B. methylotrophicum on C0 gas: ............................................. 113

Figure 25. Schematic diagram of continuous, cell recycle
CO fermentation system: ...................................................... 1 18

Figure 26. Molar product ratios as a function of pH during continuous
cell recycle fermentation: ...................................................... 125

Figure 27. Product concentrations during continuous cell recycle
CO fermentation at a pH of 5.75: ............................................. 131

Figure 28. Product concentrations during continuous cell recycle
CO fermentation at a pH of 5.5: ............................................. 132

Figure 29. Theoretical and measured washout curves for butyrate
during cell recycle fermentation at a pH of 5.5: ........................... 134

Figure 30. Known and proposed metabolic pathways for carbon and electron
ﬂow during CO metabolism in B. methylotrophicum: .................. 145

Figure 31. Molar consumption and production rates for CO and
C02 gas as a function of fermentation pH during
continuous, steady-state culture: ............................................. 150

Figure 32. Molar product ratios as a function of fermentation pH
drning steady-state, continuous culture: .................................... 150

Figure 33. Theoretical molar ATP yield via electron-transport phosphorylation and
substrate-level phosphorylation as a function of fermentation pH
during continuous culture: ...................................................... 15 2

Figure 34. Steady-state CO consumption, cell production, and optical density as
a function of pH for continuous cell recycle fermentation: .................. 15 6

Figure 35. Normalized CO consumption rates for continuous cell
recycle fermentation as a function of pH: .................................... 15 6

Figure 36. Normalized molar gas consumption and production rates
as a function of pH during steady-state, continuous
cell recycle fermentation: Case 1: ............................................. l 5 7

Figure 37. Molar product ratios as a function of pH during steady-state,
continuous cell recycle fermentation: Case 1: ........................... 15 7

Figme 38. Theoretical molar ATP yield via electron-transport phosphorylation and
substrate-level phosphorylation vs. pH during steady-state
cell recycle fermentation: Case 1: ............................................. 15 8

Figure 39. Molar electron yields as a function of fermentation pH during
steady-state cell recycle fermentation: Case 1: ........................... 159

Figure 40. Normalized molar gas consumption and production rates as a function
of pH during steady-state cell recycle fermentation: Case 2: .................. 162

Figure 41. Molar product ratios as a function of pH during steady-state
cell recycle fermentation: Case 2: ............................................. 162

Figure 42. Theoretical molar ATP yield via elecuon-transport phosphorylation and
substrate-level phosphorylation vs. pH during steady-state
cell recycle fermentation: Case 2: ............................................. 16 3

Figure 43. Molar electron yields as a function of fermentation pH dining
steady-state cell recycle fermentation: Case 2: ........................... 165

Figure 44. Electron-transport phosphorylation and total ATP yields as a

function of ET? H+IATP ratio during steady state, continuous
cell recycle fermentation: ...................................................... 1 67

Figure 45. Molar CO consumption for Cases 2-5 as a function of pH
during steady-state cell recycle fermentation: ........................... 169

Figure 46. Molar product ratios for Cases 2-5 as a function of pH
dtning steady-state cell recycle fermentation: ........................... 169

Figure 47. Theoretical molar ATP yields for Cases 2-5 as a function of pH _
during steady-state cell recycle fermentation: ........................... 170

Figure 48. Theoretical molar electron yields for Cases 2-5 as a function of pH
timing steady-state cell recycle fermentation: ........................... 170

xi

Figure 49. Normalized molar gas consumption and production
rates as a function of time dming unsteady-state
cell recycle fermentation at a pH of 5.5: .................................... 174

Figure 50. Molar yield of acetate during unsteady-state cell
recycle fermentation at a pH of 5.5: ............................................. l7 5

Figtne 51. Molar yield of butyrate during unsteady—state cell
recycle fermentation at a pH of 5.5: ............................................. 17 5

Figure 52. Theoretical molar ATP yield via electron-transport phosphorylation and

substrate-level phosphorylation during unsteady-state cell recycle
fermentation at a pH of 5.5: ...................................................... 176

xii

INTRODUCTION

The potential advantages of bioprocessing compared to catalytic processing of coal-
derived synthesis gas stem primarily from the speciﬁcity of biological conversion systems
and the associated low temperature and low pressure operating conditions. One of the
major drawbacks with using coal-derived synthesis gas as a feedstock for chemical
catalysis is the presence of sulfur gases such as hydrogen sulﬁde and carbonyl sulﬁde,
both of which are potent catalyst poisons. Removal of these contaminants is an energy-
intensive process that can add signiﬁcantly to the ﬁnal product cost, particularly for coals
with a high sulfur content. In biological processing, the conversion is conducted in an
aqueous media which can reduce direct contact of microorganisms to harmful sulﬁde
species. By employing sulfur-tolerant microorganisms, this effect can be further
minimized, and thus stringent and costly synthesis gas puriﬁcation steps can be avoided.
Furthermore, since most biological reactions occur at relatively low pressures and
temperatures, and exhibit a high degree of selectivity and efﬁciency, the cost of operation
during gas conversion can potentially be much less expensive for a biological process.
These key advantages distinguish the biological synthesis gas conversion concept and
oﬁ'er a unique niche for production of oxygenated fuels or chemicals via fermentation of
coal-derived gases.

The inherent disadvantages with biological systems are mainly slow reaction rates and
difﬁculties associated with maintenance of microbial populations in desired growth and
metabolic states. Therefore, any bioconversion scheme for synthesis gas must be designed
for maximization of reaction rates and for maintaining long-term, stable cultures. Such a
potential process scheme has been proposed by the Michigan Biotechnology Institute
(MBI). Brieﬂy, the concept employs a two-stage fermentation for production of liquid

1

fuels such as butanol and ethanol using anaerobic bacteria. The ﬁrst stage is an acidogenic
conversion where the CO present in the synthesis gas, with or without methanol, is
converted to a mixture of primarily butyric and acetic acids. The hydrogen atoms required
for this reaction come from water. The acids, along with the hydrogen gas, are then
converted in a second stage fermentation to butanol, ethanol and acetone. Cell recycling is
used in each stage to maintain high cell densities and thus high productivity. An energy-
efﬁcient distillation unit is proposed for efﬁcient, cost-effective separation of the dilute
alcohols. The ﬁrst stage CO conversion is expected to be rate-limiting and requires an
organism capable of both butyrate and acetate production directly from C0 gas. Such an
organism is Butyribacrerium methylotrophicum, an anaerobic, sulﬁde-tolerant bacterium
with the ability to metabolize a wide range of l-carbon compounds, including CO and
methanol.

The focus of this dissertation was to investigate the potential of B. methylotrophicum
for use in the MBI bioconversion scheme. A major goal of this investigation included the
design and operation of a cell recycle fermentation system for use with a gas phase
substrate. Determining the regulatory effect of fermentation pH on C0 metabolism in
studies with the cell recycle system was a key objective of those experiments. Results from
those fermentations raised several questions which were best addressed in batch metabolic
regulation studies, such as product tolerance and acid uptake phenomena. Determining the
effects of end-products and sulﬁde on C0 metabolism was a another objective. Results
from the metabolic and fermentation studies indicated that the cellular energetics could
potentially be as signiﬁcant a factor as fermentation pH in regulating the ﬂow of carbon
and electrons dming CO metabolism. To this end, a metabolic model was developed to
address the energetic hypothesis and attempt to elucidate a probable mechanism for the pH
effect, based on the results obtained in the cell recycle systems. Speciﬁcally, the key
objectives were:

1. Increase CO fermentation productivity and efﬂuent product concentrations.

2. Operate continuously at low fermentation pH values, where butyrate and

alcohol production are favored.

3. Regulate CO metabolism away from acetate and towards other products.

4. Develop a liquid media capable of supporting high cell densities.

5. Elucidate pathways for alcohol production from C0 in B. methylotrophicum.

6. Investigate effect of fermentation products on C0 metabolism.

7. Identify possible mechanism(s) of metabolic regulation during CO metabolism.

8. Determine the organism's tolerance for sulﬁde species that may be present in coal-

derived synthesis gas .

9. Determine the overall mechanism of the pH effect on C0 metabolism.

The overall objective of this work was thus to attempt to identify the mechanism of the pH
effect, both by experimentation and by modelling analysis.

The textual organization of this dissertation reﬂects the three major thrusts of this work:
regulation of metabolism, fermentation development, and metabolic modeling. Chapter 1 is
a scientiﬁc and technical review of the current chemical technology for synthesis gas
production, puriﬁcation, and utilization. The fundamentals of anaerobic, acetogenic
bacteria are then discussed, including a review of the most widely known organisms
capable of metabolizing the major synthesis gas components. The origin, physiology, and
metabolism of B. methylotrophicum are then reviewed, followed by a description of the
MBI process for biological conversion of coal-derived synthesis gas.

Chapter 2 is a description of the materials and techniques used during the experimental
portion of this work. Chapter 3 presents the microbiological studies conducted in
anaerobic bottle systems. These studies included the effect of all major CO fermentation
products, including acetate, butyrate, ethanol, and butanol, on C0 metabolism during
growth. The effect of sulﬁde species on C0 metabolism during cell growth is then
described. Finally, the development of a culture medium for use in the cell recycle

fermentations is presented.

Chapter 4 encompasses the cell recycle fermentation studies conducted with B.
methylotrophicum. This chapter is divided into sections based on the observed metabolic
responses to fermentation pH, mainly steady-state responses, oscillatory responses, and
cell death. These results are discussed in light of the current hypothesis for the mechanism
of the pH effect on C0 metabolism, including the possibility of energetic limitations during
CO fermentation. Experiments relating to the use of cosubstrates during CO fermentation
are then presented.

Chapter 5 presents the development and analysis of results of the metabolic model for
CO metabolism, based on the cell recycle fermentation data. The basic principles and
objectives of the model are outlined. The model is then applied to previously derived
chemostat data, the results evaluated, and the modeling approach evaluated. The steady-
state cell recycle fermentation results are then analyzed with the model. The energetics of
CO metabolism are discussed in light of the fermentation analysis. A preliminary analysis
of the unsteady-state fermentation data is then presented. A short discussion on the
limitations of the modelling method concludes this chapter. The ﬁndings from Chapters 3-
5 are summarized and discussed further in Chapter 6, with a focus on the pH effect and its
likely mechanism. Finally, the work is concluded with a statement concerning the potential
for a bioconversion process based on the presented results, and with some

recommendations on areas and aspects for further investigation.

CHAPTER 1

TECHNICAL AND SCIENTIFIC REVIEW

1.1. Synthesis Gas Chemical Technology

1111] C EC ”1.1..

The cultural and political volatility and instability of the world's major petroleum
producing regions as well as the predicted future shortages in petroleum resources has
stimulated both discussion and research on alternate forms of energy and their utilization.
Whether these issues and many others are sufﬁcient to motivate the United States to
implement policy changes towards a major energy transition remains to be seen.
Diversiﬁcation of energy resources instead of the current heavy dependence on a
petroleum based energy and fuels economy would certainly be advantageous from a
domestic standpoint. However, the competitive advantage of petroleum versus other
energy forms, based on its abundance, and based on the dependence of the current
infrastucture on its use, makes diversiﬁcation a difficult and arduous process. The ultimate
degree of diversiﬁcation obtainable depends on development of policy, technology, and
also on efficient application of the technology, to modify and refocus the current structure
of the industry. Social and environmental factors will also have an important inﬂunce on
the degree of diversiﬁcation possible. Such non-technical factors have even been described
as the major obstacles towards commercialization of new processes (Conn, 1981).
Finding substitutes for petroleum to serve as sources for fuel, energy, and chemicals
production is thus an important, difﬁcult, and far-reaching endeavor.

The use of coal as direct energy source or as a raw material for fuels and chemicals
production is an existing and increasingly attractive alternative to petroleum-based
processes. In 1990, US. coal production surpassed 1 billion tons for the ﬁrst time

(Lopkowski, 1991), and is predicted to surpass 2 billion by the year 2000 (Exxon Co.,
5

1979). Currently, coal-ﬁred plants account for 56% of domestic power generation, with
utility companies consuming approximately 80% of the total mined coal (Haggin, 1991).
Another major industry based on coal is the coking industry, in which carbonization of
coal via destructive distillation produces coke, which has a principal use as a fuel in blast
furnaces, with wide application in the steel industry. In the fuels and chemicals sector,
coal's contribution is much less evident, with the majority of fuels and chemicals
production currently arising from natural gas and petroleum utilization. The domestic
abundance of coal, its comparative cost, and a relatively large processing technology base
are some of the advantages associated with its use. Environmental issues and
transportation costs are a few of the major disadvantages.

In general, coals are ranked or categorized by their relative aromaticity and their
moisture and volatile matter content, with anthracite, semi-anthracite, low-volatile
bituminous, medium-volatile bituminous, hi gh-volatile bituminuous, sub-bituminous, and
lignite coals comprising the major classiﬁcations. These types are listed in order of
increasing moisture content and volatile matter content and decreasing aromatic content. A
high rank coal is thus one in which the ﬁxed carbon to volatile carbon ratio, with ﬁxed
carbon in the form of an aromatic ring structure, is very high. The above list is also
indicative of the overall energy content for combustion, most commonly listed in the units
of Btu/lb, where lignitic coal has the lowest caloriﬁc value and anthracitic coal the highest.
It follows that lignitic coal is typically cheaper in comparison, mainly because the price is
generally determined by the application, which for coal is mainly in its combustion for
electric power generation. Low-rank coals have thus become a focal point for synthetic
fuels and chemicals research, because of their relatively low cost.

The complexity and diversity of coal has been one of the limiting factors in
development of new technology for speciﬁc applications. The diversity of structure stems
primarily from the diversity of carbonaceous plant life that was fossilized, the geologic

environment in which the fossilization occurred, and on the age of the coal. All of these

factors are reﬂected in the composition of both the organic and inorganic structure.
Although the ﬁne organic structure of coal is a topic still under debate and investigation, it
can generally be described with an empirical formula containing carbon, hydrogen,
oxygen, nitrogen, and sulfur such as CHO,70(),6N0,280.1 for bituminous coal (Datta and
Andrews, 1991). The inorganic fraction is typically in the form of discrete mineral
inclusions. Contained in both the organic and inorganic fractions are sulfur compounds,
present primarily as inorganic iron pyrite or covalently bound organic thiols,
sulﬁdes/disulﬁdes, and mercaptans (Young and Finnerty, 1991). The overall sulfur
content of coal is one of the key characteristics in determining its value and utility, since
these sulfur compounds will result in the formation of sulfur oxides during combustion,
which will have a direct impact on air quality and emissions during processing.

These factors have been one of the key inﬂuences in the deﬁnition and implementation
of the US. Department of Energy's clean-coal program, which has been focused on
increasing the uses and applications of coal in the national energy strategy. Although this
effort has been primarily based on restructuring and redesigning existing technology for
using coal in electric power generation, other areas are also under investigation. One area
which is very active is the development of lower rank coal and coal-based processes for
synthetic fuels and chemicals/feedstocks production. Most of this developmental research
is in optimization of technology for design of commercial facilities and plants (Bassler and
Sperhac, 1986). These routes involve coal liquefaction technology.

1 l 2 11' I 1. I . E .

Coal utilization and processing for synthetic fuels and chemicals production can be
categorized under 4 generic coal liquefaction methods, which are pyrolysis, direct
hydroliquefaction, solvent extraction, and indirect synthesis. If synthesis gas is produced
as an intermediate, then indirect synthesis is referred to as indirect liquefaction, while all
the others are referred to as direct liquefaction. Hydroliquefaction is generally accepted to

be the preferred direct liquefaction route, and is a technology that has its roots in coal

pyrolysis. Solvent extraction is not currently under serious development, particularly on
the pilot and commercial scales. Indirect liquefaction has been shown to be a commercially
viable technology, albeit under speciﬁc circumstances, as discussed below.

Coal pyrolysis, also known as destructive distillation, is a process by which the coal is
thermally decomposed to form a mixture of gases, liquids, tars, char, and coke. The
principle mechanism is ﬁrst removal of the volatile fraction, and then generation of organic
free radicals, which subsequently react with the solid structure to form a mixture of
volatile liquids and gases. A carbonaceous char is the resulting solid structure, which is
typically reacted or gasiﬁed further. A major limitation is the availability of hydrogen,
which during pyrolysis comes from the coal itself. In the abscence of available hydrogen,
polymerization and aromatization of the reactive free redical coal fragments occurs,
resulting in excessive coke formation (Mills, 1976). The regulation of destructive coal
distillation is thus dependent on the relative rates of free radical generation and on the rates
of hydrogen tranfer to the free radical reactive intermediates; balanced rates will favor
volatile hydrocarbon formation, low rates of hydrogen donation and transfer will favor
coke formation. Some pilot-scale processes which demonstrate pyrolysis technology are
the COED process, the TOSCOAL process, and the Occidental (Garrett Coal Pyrolysis)
process (Kalfadelis and Shaw, 1980).

In direct hydroliquefaction, which is also called coal hydrogenation, coal is thermally
liquiﬁed in the presence of catalyst and high pressure hydrogen. The basic mechanism is
the same as for pyrolysis, only in hydroliquefaction, external hydrogen helps to minimize
the reactive polymerization of the free radicals, which thus minimizes coke and char
formation. This hydrogen can come from molecular hydrogen, or more typically from a
hydrogen-donating solvent added to the reaction mixture. Balancing the reaction
temperature, which as it is increased promotes free radical formation, with the reaction
pressure, which as it is increased allows for higher hydrogen availability until catalyst

saturation, is a key issue in these processes. Some representative technologies are the

Synthoil process, the H-Coal process, and the Exxon donor-solvent coal liquefaction
process (Mangold et al., 1982).

Indirect liquefaction is a technology based on intermediate formation of synthesis gas
by coal gasiﬁcation, and then further catalytic reaction to produce the desired liquid
products. Coal is gasiﬁcd by thermal reaction in the presence of oxygen and water to
produce carbon monoxide (CO) and hydrogen (H2) gases. These gases can then serve as
precursors for several different reaction chemistries, either direct or indirect themselves,
including Fischer-Tropsch synthesis, Roelen chemistry, which is also known as oxo-
processes or hydroformylation, and many indirect methanol chemistry routes. These
technologies have a long history, are currently in use commercially, and are discussed in
more detail in a subsequent section.

“3513’ 151.521.

Indirect coal liquefaction requires gasiﬁcation of coal to form synthesis gas, primarily a
mixttne of CO and H2 gases. Although synthesis gas for fuels and chemicals production
is presently produced primarily by catalytic steam reformation of natural gas or by partial
oxidation of heavy oil fractions (Heubler and Janka, 1980), coal gasiﬁcation to form
synthesis gas has a future potential because of the relative abundance, availability, and cost
of coal. Coal gasiﬁcation in the US. between the turn of the century and until the 1950's
was an attractive technology, with extensive production of gaseous fuels occuring in
plants which operated generally at low capacity and low efﬁciency (Spencer et al., 1982).
This technology, however, was superseded by the availability of low cost natural gas and
petroleum after the Second World War.

The revitalization of coal gasiﬁcation technology depends on the redesigning of old
methods to produce more economically and environmentally attractive processes. With
gasiﬁcation becoming a more prominant step in electric power generation from coal, the
electric power industry has served to develop gasiﬁcation technology to a relatively high

level, technology which can directly translate to synthetic fuels and chemical production

10

via indirect liquefaction. Typically, synthesis gas is produced from coal by steam
gasiﬁcation, where coal, water, and oxygen (02) are combined at high temperature and
pressure under conditions promoting an O2 limited system. The resulting synthesis gas is
primarily a mixture of CO and H2, with carbon dioxide (CO2), methane (CH4), hydrogen
sulﬁde (H28), and H20 being the remaining components. The composition of the
synthesis gas depends mainly on the type of gasiﬁer and the resulting reaction conditions.
The conditions directly affect 5 main reactions occuring in the gasiﬁer (Mangold et al.,

1982):

C + 0.502 ---> CO (1)
C + C02 ---> 2C0 (2)
C0 + H20 ---> CO2 + H2 (3)
C + 2H2 ---> CH4 (4)
C0 + 3H2 ---> CH4 + H20 (5)

The relative rates of these reactions will determine the exit synthesis gas composition.
Under oxygen limited conditions, reaction (1) becomes rate limiting thus making
incomplete or partial combustion of coal a fundamental step. Reaction (3) is often referred
to as the water-gas shift reaction, and is the main regulatory reaction in determining the
ﬁnal CO:H2 ratio. This reaction also favors H2 at lower temperatures. Reaction (5) is
what has been referred to above as steam reformation of natural gas, and is a reaction
which regulates CH4 formation in the presence of steam (H20). The ﬁnal synthesis gas
composition desired depends on the application; for substitute natural gas (SNG)
production, one wants to maximize CH4 and H2 formation and minimize C0 formation.
For electric power generation, one wants to maximize the CO:H2 ratio, and for chemical
and fuels production, a speciﬁc CO:H2 based on the stoichiometry based on the product-
forming reaction is usually desired. These objectives are most commonly achieved by the
choice of gasiﬁer system, whether it be a ﬁxed- or moving-bed, a ﬂuidized-bed, or an
entrained ﬂow gasiﬁer (Wilson et al, 1988).

11

The design and performance of the main gasiﬁer types are based primarily on the
physical and mechanical aspects of contacting the coal with the 02 and H20 .
Fixed/moving-bed gasiﬁers operate with countercurrent ﬂow of coal and reacting gases,
generally operate below 10000F, and exhibit a very non-uniform temperature proﬁle in the
reactor. These systems generally cannot react the coal ﬁnes, and produce heavy by-
product tars (Spencer et al., 1982). Relatively high solid residence times are required, and
the mechanical strength of the coal must be high (V erma, 1978a). Entrained—bed gasiﬁers
use cocurrent ﬂow of coal and reactant gases, generally operate between 2300°F and
27009F, and exhibit a very uniform temperature proﬁle. Reaction of coal and carbon ﬁnes
in these systems is possible. Methane and tar formation are not observed because of the
high temperature of operation. Fluidized-bed systems exhibit rapid and complete gas-
solids mixing, operate between 1000°F and 17000F, and exhibit a relatively uniform
temperatme proﬁle. The characteristics described above are critical in determining the ﬁnal
synthesis gas composition (Wilson et al., 1988).

The application of these reactor types and their development has resulted in a so-called
second generation of coal-gasiﬁcation systems on a world-wide basis. Only a few types
and modiﬁcations of these have been commercially proven, such as the Wellman-Galusha,
Winkler, Koppers-Totzek, and Lurgi type gasiﬁers, but many others are at the pilot scale
of development and testing. Most of the systems were designed initially for production of
gases in a particular heating value range, but current technology now allows a certain
degree of versatility in this regard. All of these types were invented in Germany over 50
years ago.

The Wellman-Galusha gasiﬁer is a ﬁxed-bed conﬁguration designed speciﬁcally for the
production of low-Btu gases. The reactant gas stream is a mixture of steam and air, and
thus nitrogen is typically a signiﬁcant portion of the resulting synthesis gas. Operating

pressures are atmospheric or slightly above atmospheric pressure. The temperatures of

12

operation are in the range of 1000-12000F (Verma, 1978a). A major operating advantage
of this design is its ability to process all types and ranks of coal.

The Winkler gasiﬁer is a ﬂuidized-bed type unit in which crushed coal is fed by screw
feeder to the bottom of a fuel bed, where the gas stream is introduced and the reacting
mixture ﬂuidized. The ﬂuidizing gas used is usually a mixture of steam and air or 02. The
unit typically operates at 1500-18509F and at atmospheric pressure (V erma, 1978a).
Conversion of carbon in the coal is relatively low. The Winkler process was one of the
ﬁrst applications of ﬂuidized-bed technology, but has been transcended by competing
technology in recent times. However, a version of this type of gasiﬁer, called the High
Temperature Winkler (HTW) or Rheinbraun process is under development at the pilot and
demonstration plant scales. This version of the conventional Winkler technology uses a
high pressure gasiﬁcation system nearing 10 atrn (Simbeck, et al., 1982).

The Koppers-Totzek (K-T) gasiﬁer is an entrained-bed type where dried and ﬁnely
crushed coal is conveyed into a mixing nozzle, where it is contacted by a jet of steam and
O2 and entrained into the gasiﬁer. Reaction pressures are nominally above atmospheric,
with reaction temperatures nearing 2600°F. Carbon conversion is relatively high. The K-T‘
process is one of the most recent gasiﬁers to be proven commercially, with the primary
application in the production of synthesis gas for ammonia formation (Simbeck et al.,
1982). There are several offshoots of the K-T process under development at the pilot and
demonstration scales, most notably of which is the Shell-Koppers process, which uses a
pressurized version of the K-T gasiﬁer (Simbeck et al., 1982).

The Lurgi gasiﬁer is perhaps the most successful of the commercialized gasiﬁer
technologies. Its conﬁguration is an 02/steam blown, high pressure, moving-bed.
Operating temperatures are in the range of 1150-14000F. The countercurrent ﬂow
conﬁguration results in various reaction regimes in the gasiﬁer, with zones of combustion,
gasiﬁcation, carbonization, and drying occurring. The primary examples of commercial

application are in the South African Coal, Oil, and Gas Corporation, or SASOL, where

13

Lurgi gasiﬁers are used to produce synthesis gas for subsequent Fischer-Tropsch
processing and liquid fuels production. Lurgi gasiﬁers are also employed at the Great
Plains Coal Gasiﬁcation Plant, in North Dakota (Imler, 1985). A version of the Lurgi
which reduces steam requirements by slagging the ash produced in the gasiﬁer, called the
BGC/Lurgi process, is in the demonstration scale of development (Simbeck et al., 1982).
Current developments in gasiﬁer technology are coming mostly in the areas of reduced
utilities requirements and improved capacities. A wide range of synthesis gas
compositions are obtainable as shown in Table 1; these compositions are not only
dependent on the gasiﬁer design and operating conditions but also on the type of coal

used.

Table 1. Synthesis gas compositions from commercial and near-commercial coal

 

 

gasiﬁcation processes.
Gas Composition (vol %)

Process 00 H2 C02 H2O CH4 H28 + COS
Wellman-Galushaa 22.8 22.8 11.4 0.0 0.0 (43.0 N2)
Winklera 36.0 42.0 17.7 0.0 3.2 0.1
Koppers-Totzeka 52.5 39.0 10.0 0.0 0.0 0.5
Ltn'gia 21.0 41.0 27.5 0.0 8.8 -
Shell-Koppersb 64.0 31.6 0.8 1.5 0.0 1.4
BGC Lurgic * 55.5 28.3 5.1 0.0 6.3 3.0
Texacod 54.0 34.0 11.0 0.0 0.01 0.3
Underground Gasiﬁcatione 19.6 31.1 33.0 - 16.3 -

 

a From Verma, 1978a.
b From McCullough et al., 1982.

C From Development Status of Key Emerging Foreign Gasiﬁcation Systems, TRW
Report. Sept, 1980.

d From Schafer et al., 1988.
c From Bloomstran and Davis, 1984.

Several large energy corporations, including Texaco (Kolaian and Schlinger, 1982) and

Shell (McCullough, 1982), are currently developing demonstration scale variations of the

14

basic types of gasiﬁers listed above (Spencer et al., 1989; Schafer et al., 1988). However,
gasiﬁer development is not the only area of technology advancement for coal gasiﬁcation.
For speciﬁc applications, mainly in the production of low-Btu gases, underground coal
gasiﬁcation in situ, for deep coal seams where mining is either impractical or impossible,
has undergone favorable ﬁeld tests in recent years (Hammesfahr and Winter, 1984;
Bloomstran and Davis, 1984). This technology, ﬁrst developed in the Soviet Union over
60 years ago, employs air or 02 and steam injection through bore holes drilled into the
coal seam. The drilling process and the gas ﬂow properties in coal strata are the major
obstacles to commercial development (V errna, 1978b).

1 1 l E '6 . [S l . Ci

The production of synthesis gas from coal is a technically complex process. The
complicated structure and composition of coal results in a variety of contaminant species
being formed during gasiﬁcation, most of which are sulfur and nitrogen compounds.
Formation of C02 in conventional gasiﬁers, as shown previously in Table 1, can account
for anywhere between 1-30% by volume of the synthesis gas mixture. These impurities,
or contaminants, must be removed prior to catalytic synthesis, because of catalyst
poisoning limitations. There are also many corrosion and environmental/safety issues
involved in sulfur and acid gas processing. Separation and removal is not the only
concern, however. Once the synthesis gas has been puriﬁed, the contaminant sneams
must be recovered, recycled, or disposed of in some way. The magnitude of these
operations is enormous when the processing rate of most commercial gasiﬁers is realized.

The presence of sulfur compounds in coal can vary widely, with the sulfur content
range between 0.5-11% by weight for US. coals (Young and Finnerty, 1991). Inorganic
sulfur pyrites (Fe82) are generally removed in density—dependent pretreatment processes
prior to gasiﬁcation, since these compounds are generally present as distinct mineral
inclusions. It is thus the organic sulfur in coal that contributes during gasiﬁcation to the

formation of contaminating sulfur species. The reducing conditions of the typical

15

gasiﬁcation process result in most of the sulfur in coal ending up as hydrogen sulﬁde
(H28) or carbonyl sulﬁde (COS), with the vast majority as H28. As shown previously in
Table 1, these compounds can typically account for between 0-3% by volume of synthesis
gas mixtures, with the ﬁnal concentration highly dependent on the composition of the coal
gasiﬁed. Other sulfur compounds that are often observed in much lower amounts are
carbon disulﬁde (C82), thiophenes, mercaptans, and several alkylated sulﬁde Species
(Cumare et al., 1984).

Nitrogen contaminants in synthesis gas can be of particular concern if the gasiﬁcation
process is using air and steam instead of pure O2 and steam. The nitrogen species from
operation with air are typically NOx type species, including N20, NO and N02. The use
of 02 instead of air will generally limit formation of these oxidized nitrogen compounds
(Cumare et al., 1984). Organic nitrogen contained in the coal will be liberated during
gasiﬁcation in the form of ammonia (NH3), hydrogen cyanide (HCN), and elemental
nitrogen (N2). Since the nitrogen content of coal, on the average, is between 1-2% by
weight, these impurities are less prevelant than sulfur impurities.

The necessity and stringency of contaminant removal for synthesis gas processing is
determined primarily by catalyst poisoning issues, the vast majority of which concern the
concentration of sulﬁdes in the reactant synthesis gas. As detailed in the next section, both
production of substitute natural gas (SNG) and conventional Fischer—Tropsch chemistry
are catalyst-dependent. These processes have different sulfur tolerance levels based on the
catalyst composition. Catalysts which have been determined to be active in CO
hydrogenations such as those observed during Fischer-tropsch synthesis include most of
the Type VIII metals, namely iron, cobalt, nickel, ruthenium, platinum, and also rhodium.
0f the above metals used in synthesis gas processing, iron is the most common and the
best characterized.

Speciﬁcally, the poisoning of Type VIII metal catalysts is due primarily to the reactivity

of the two unshared electron pairs present in a molecule of H28, which contribute to

16

stability during metal-poison interactions (Lee, 1985). The strength of these interactions is
explained in part by a high heat of adsorption for sulﬁde on most of the metals listed
above. Besides the issue of active site poisoning, there is evidence from work with
platinum for decreased CO binding energy with increased sulfur coverage, possibly due to
electronic interactions between the adsorbed sulfur and the adsorbing CO (Bonzel and Ku,
1973). This same study, using ﬂash desorption techniques, determined that sulfur and CO
compete for the same active sites on platinum catalysts. The negative effect due to sulfur
poisoning on the binding energy of CO is partially a consequence of a weakening of the
electronic density in sites adjacent to that of the adsorbed sulfur molecule, a result
stemming from its highly electronegative nature (Oudar, 1980). In the case of surface
reactions characterized by a Langmuir-Hinshelwood mechanism, the presence of
chemisorbed sulfur is detrimental, a likely effect of decreased active species mobility
(Hegedus and McCabe, 1984).

The actual permissible sulfur concentrations in a typical synthesis gas feedstock are
quite low. For iron at 400°C, 1.6 ppm is the tolerance limit, and for nickel at 550°C, the
limit is 2.5 ppm (Rostrup-Nielsen, 1971). Current industrial standards aim at an order of
magnitude less than these values, to maintain quality control and long catalyst life (Dry,
1981). These low levels are obtainable by employing advanced sulfm' removal technology,
as described below.

The predominant contaminant species H28 and CO2, also known as the acid gases,
account for the majority of synthesis gas separation and removal technology, with removal
of C08 and C82 occuring simultaneously. H28 removal is necessary to prevent corrosion
of process equipment, to protect downstream catalytic reactors from poisoning, and to
maintain saftey and environmental quality in the plant. Removal of C02 is necessary
during production of high-Btu synthesis and fuel gases. There are two basic commercial
technologies for removal of these compounds, either solvent absorption or chemical

conversion. The choice between these system depends heavily on the partial pressures of

l7

acid gases in the synthesis gas stream. Solvent absorption processes can be further
separated by the nature of the solvent used, with chemical, physical, and hybrid solvent
absorption strategies all being commercially available. The chemical solvents allow
chemisorbtion with acid gas compounds. These solvents usually have high heats of
reaction with acid gases, and possess high heats of absorption for the same compounds
(Eickmeyer and Gangriwala, 1981). The physical solvent processes exhibit a much
different equilibrium curve than chemical solvents, and are useful at high pressures and
higher acid gas partial pressures. Hybrid solvent systems are mith of both chemical and
physical solvents and attempt to combine the best characteristics of each, and are used
when very speciﬁc absorption characteristics are required.

The majority of chemical absorption systems are amine-based solutions which are
weakly alkaline. The alkanoamines such as monoethanolamine (MBA), diethanolamine
(DEA) are the most common, with DEA usually preferred because of the negative effect of
C08 on MEA (Cumare, 1984). The use of amine-based solvent systems is limited by the
temperature-equilibrium curve for the acid gas components; increasing temperature results
in a sharp increase in the CO2 back pressure for these chemisorption systems (Eickmeyer
and Gangriwala, 1981). The heat tranfer issues during removal are thus paramount when
employing amine-based systems. Other possible solvents include diglycoamine (DGA), a
primary amine which is similar to MBA but more soluble in water. The secondary amine
diisopropanolamine (DIPA) is speciﬁc for H28 removal, and is also used in various
hybrid solvents. The tertiary amine methyldiethanolamine (MDEA) is also selective for
H28 over C02, and is used in acid gas separations. In general, one positive feature of all
amine-based chemical absorption systems is the low solubility of hydrocarbons in the
aqueous amine-water solvents (Cumare, 1984). When the operating pressure and the
partial pressures of the acid gas contaminants are adequately high, chemical solvents based
on aqueous hot potassium carbonate solutions are preferable, because of the lower heats of

absorption for CO2 and thus lower utility costs (Eickmeyer and Gangriwala, 1981). When

18

a catalyst is added, called promoted hot carbonate solutions, the rate of CO2 absorption is
increased even more. These systems also exhibit low solubilities for hydrocarbon
absorption. Commercial hot carbonate processes include the Catacarb, Flexsorb HP, and
Benﬁeld processes.

Absorption of acid gas constituents in physical solvents is applicable at higher
pressures and higher partial pressures of acid gases. High acid gas loading in physical
solvents allows higher acid gas removal capacities and lower solvent circulation rates
(Cumare et al., 1984). One major drawback of using physical solvents is the absorption of
heavier hydrocarbons compared to chemical solvents (Cumare et al., 1984). The use of
physical solvents is often determined by the process heat requirements and availability of
process heat for regeneration of the solvent, mainly because steam stripping is required for
removal of H28 (Eickmeyer and Gangriwala, 1981). Representative commercial physical
solvent processes include Selexol, Rectisol, Purisol, Sepasolv MPE, and Catasol, the
selection of which can depend on which acid gas component is more critical for removal,
athough all of the above will absorb both H28 and CO2 to a certain degree (Cumare et al.,
1984).

Hybrid solvents are a combination of chemical and physical solvent systems and exhibit
properties of both. Shell's Sulfmol is perhaps the best known, which is a combination of
DIPA and the physical solvent Sulfolane. Hybrid solvents generally have a lower water
content than amine-based chemical solvents and thus allow neatment of acid gases at
higher partial pressures (Eickmeyer and Gangriwala, 1981). Other commercial examples
are Amisol and Flexsorb PS.

There are several puriﬁcation processes that are currently in the development and testing
phases of research to explore more innovative and economical puriﬁcation technologies.
The use of pressure-swing adsorption phenomena to separate individual coal gas
components is one such approach. Application of speciﬁc ion-exchange membranes is

anorher. Dissolving chemical coordination complexes in an organic solvent for speciﬁc

19

synthesis gas component separation is also a current approach. One of the most intriguing

methods is the use of speciﬁc gas separation membranes for CO and H2 puriﬁcation and
separation, with application in controlling CO:H2 ratios for synthesis gas processing.
These membranes can operate at very high pressure, can employ cryogenic, catalytic, and
pressure swing adsorption phenomena, and are relatively simple to operate (Spillman and
Sherman, 1990). The preliminary economics of these membrane-based separation systems
are favorable compared to conventional gas separation schemes (Spillman, 1989). In
general, all of these novel ideas are still in the exploratory phases of research and
development, but can potentially offer substantially reduced operating and capital costs
compared to the present separation and puriﬁcation technologies available (Wilson et al.,
1988).

Chemical conversion processes to remove acid gases and other synthesis gas
contaminants are designed either for direct removal or for sulfur recovery following direct
removal by solvent-based systems. The use of chemical conversion processes for direct
sulfur removal requires synthesis or raw gas streams with low H28 partial pressures
where high H28 selectivity is required (Cumare et al., 1984). When these conditions are
present, direct chemical conversion is possible by using commercially available processes
such as the Stretford process, and the Takahax process. Generally, these types of chemical
conversion require less than 10% 1128 in the gas stream and are used when CO2 removal
is not necessary or warranted. The Stretford process is a catalytic vanadium oxidation of
hydrosulﬁde ion, with very high puriﬁcation capability. The Takahax process is similar in
design but uses naphtoquinonc (Cumare et al., 1984). When necessary, speciﬁc catalytic
processes for conversion of C08 gas to H28 gas can be employed, such as the CKA COS
hydrolysis reaction catalyst (Udengarrd and Berzins, 1984).

The typical process ﬂow for synthesis gas puriﬁcation is thus dependent on the
composition of H28, CO2, and C08 in the raw synthesis gas, and upon whether or not

higher-Btu fuel gases are desired. Low sulﬁde containing raw gases with no requirement

20

for C02 removal can be treated directly by chemical conversion and the sulﬁde species
removed. Synthesis gas streams requiring C02 and H28 removal are ﬁrst treated using
either chemical or physical solvent absorption systems, typically physical solvent systems
for moderate to high H28 contents, and chemical solvent systems for moderate to low
H28 contents. The resulting acid gas stream, which is a mixture primarily of H28, CO2,
and other sulﬁde species, is then treated to selectively remove the CO2 from the H28, a
degree of removal which is ultimately dependent on the type of sulftn' recovery technology
employed downstream. Figure 1 outlines a typical synthesis gas puriﬁcation process ﬂow
concept.

Sulfur recovery technology is based on the proportion of H28 in the separated acid gas
stream. The original chemistry for coverting H28 to elemental sulfur was invented by Carl
Friedrich Claus, an English scientist working in the late nineteenth century. In the late
1930's, German scientists redesigned and improved the Clans process and deveIOped
what is now known as the modiﬁed Claus process. It is estimated that the Claus process
accounts for approximately 90-95% of the world's recovered sulfur (Goar, 1986).
Typically, the Claus process can recover 95-97% of the feed H28. However, recent
advances and new processes and modiﬁcations have increased the obtainable recovery to
over 99% (Goar, 1986).

The Claus and modiﬁed Claus process are basically catalytic combustions of H28 in the
presence of O2, governed by the two reactions,

H28 + 1.502 ---> 802 + H20 (1)
H28 + 802 ---> 1.582 + 2H2O (2)
which are called the Burns and Claus reactions, for reactions 1 and 2, respectively. The
overall reaction taking place is thus,
H28 + 1.502 ---> 1.582 + H2O.
Numerous side reactions are possible and do occur during the combustion reactions,

including formation of carbonyl sulﬁdes, carbon disulﬁdes, and various oxidixed sulfur

21

 

 

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22

species. The typical process scheme is to have several Claus reactors in series for
increasing recovery.

The limitations of the Claus technology are in the minimum H28 concentration required
for efﬁcient operation, which was previously 30% but has been steadily improved over the
last 50 years (Cumare et al., 1984). These improvements to the Claus process have
included preheating of both air and fuel, sulfur recycling and burning, and integration with
selective solvent absorbtion units to increase the fraction of H28. Such modiﬁcations have
allowed recovery of sulfur from gas streams between 5-30% H28 (Cumare, 1984). For
acid gas streams less than 5%, a more recent technology called the Selectox process is
typically used, in which the entire recovery process is catalytic and no combustion of H28
to 802 is done (Goar, 1986).

After a majority of the sulfur is recovered by the Claus or Selectox processes, the
remaining gases are either vented to the atmosphere or proceed downstream to tail gas
clean up processes, depending on the sulfur composition of the exit gas. Tail gas clean up
is typically a wet scrubbing or a dry bed process, where the remaining sulfur compounds
are either oxidized or reduced and subsequently reacted again. Therefore, tail gas clean up
technology is usually a reﬁned extension of conventional modiﬁed Claus sulfur recovery
technology (Cumare et al., 1984).

1155 l .5 11.1..

The uses and potential applications of synthesis gas are widespread. Electric power
generation, substitute natural gas production, methanol synthesis, and a host of fuels, fuel
additives, fuel oils, and chemicals are obtainable from l-carbon chemistry and synthesis
reactions from C0 and H2. These potential routes are outlined in Figure 2. Many of these
technologies have long, proven histories, others are new, as yet commercially unproven
processes, and still others are modiﬁcations to existing technologies. Whatever the
application, the employment and economics of coal-derived synthesis gas technologies are

ultimately dependent on the limitations of coal as an energy source, as discussed earlier.

23

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24

The electric utility industry has by far been the major consumer of coal in the US, but
ﬂuctuating natural gas and petroleum prices, restrictive environmental regulations in the
form of compliance issues for emissions reduction and control, and aging plant hardware
problems have been pushing the industry towards a ﬁnancial crisis (Spencer et al., 1982;
Haggin, 1991). Consequently, a restructuring of the industry has been taking place to
streamline the economics of electricity generation, in order to stay competitive while
meeting environmental requirements. The major effort in this restructuring has been a shift
from conventional coal-ﬁred plants with oxidized sulfur scrubbers to integrated coal
gasiﬁcation combined-cycle (IGCC) technology.

IGCC technology is generally in the site-testing and demonstration phases of
commercial development, and the outlook is favorable, particularly because of the
differences in sulfur emissions problems arising from plant operation (Spencer et al.,
1989). The basic process design is to gasify the coal to form a CO-rich synthesis gas,
from which haet is removed to produce high-pressure, saturated steam from process
water. The steam is then superheated in heat recovery steam generators, from which it then
goes to steam turbines to produce electricity. The cooled synthesis gas, which has a high
sensible heat content, passes through sulfur and nitrogen puriﬁcation steps, as described
in the previous section. The clean synthesis gas is then expanded and combusted in gas
turbines, producing electricity and providing more heat for low-quality steam generation.
Technical improvements in both gasiﬁer and gas turbine systems have enhanced the
potential for IGCC electric power generating systems, although the gas cleaning and
environmental emissions compliance are still a signiﬁcant portion of both the capital and
total cost (Haggin, 1991).

An interesting design modiﬁcation to the basic IGCC plant design, and one which
reﬂects the versatility of C1 chemistry, is cogeneratin g or coproducing power plants which
produce both electricity and medium to high-Btu gas or chemicals. These designs split the

synthesis gas stream entering the gas turbines for processing to SNG, for sale "over the

25

fence" as medium-Btu fuel gas, or for catalytic, onsite production of methanol (Spencer et
al., 1982). The methanol can then be used as a fuel on site or sold as a commodity. In
general, these potential modiﬁcations have not yet been adequately tested, but are
indicative of the overall ﬂexibility and futme for IGCC power plants.

The Btu content of coal-derived synthesis gas is dependent on the gasiﬁcation process,
which is usually dictated by the application. One route for energy generation from coal is
in the production of high-Btu SNG from synthesis gas. This process involves ﬁrst
gasiﬁcation and then methanation, and requires synthesis gas with a low CO:H2 ratio.
Such a ratio is controlled by regulation of the water- gas shift reaction,

C0 + H2O ---> C02 + H2
both in the gasiﬁer and in downstream shift conversion reactors. Typically, a 3:1 H2 to
CO ratio is required prior to methanation, and the ratio is controlled by splitting the
synthesis gas stream and remixing shifted and unshifted streams to obtain the ﬁnal
proportion. One of the current topics of gasiﬁer development for this technology is in the
design of coal gasiﬁcation systems which are high in their CH4 and H2 compositions
relative to CO, in order to minimize the requirements for a methanation step. The synthesis
gas stream, once the required CO:H2 ratio has been obtained, is then cooled, pmiﬁed, and
passed to the methanation units, where the exothermic reaction,

C0 + 3H2 ---> CH4 + H20
is mediated using a nickel catalyst to produce SNG (V erma, 1978a). The heat of reaction
for this methanation step is used to generate process steam, which is subsequently super-
heated and used to drive turbines for compressor duty.

The best commercial example for this technology is the Great Plains Coal Gasiﬁcation
Plant, in Mercer County, North Dakota, which is designed to produce pipeline quality
SNG from lignite coal. This plant employs 12 Lurgi Mark IV gasiﬁers and uses a Rectisol
unit for sulfur and C02 puriﬁcation (Imler, 1985). An adjacent coal-ﬁred power plant

provides the electricity for the gasiﬁcation and puriﬁcation operations. One of the largest

26

grassroots facilities in the US, the Great Plains Coal Gasiﬁcation Project was the nations
ﬁrst commercial synthetic fuels plant based on coal conversion technology (Imler, 1985).
For producing liquid fuels and chemicals from synthesis gas, the available technology
can be separated into direct and indirect synthesis routes. The most notable direct synthesis
routes are the Fischer-Tropsch synthesis and the oxonation (hydroformylation) or Roelen

chemistry synthesis. Indirect routes are typically through methanol chemistry, or by using
methanol (CH3OH) plus CO/Hz synthesis schemes. Between the direct and indirect

synthesis routes, the potential applications of C1 chemistry technology based on coal-
derived synthesis gas are widespread, as outlined in Figure 3. These chemisty of these
direct and indirect synthesis routes are discussed in more detail below, including the
current commercial applications.

The Fischer-Tropsch synthesis is a catalytic conversion of CO and H2 to variety of
parafﬁnic, olef'mic, and oxygenated compounds. These routes were developed in the
1920's and 1930's by Franz Fischer and Hans Tropsch in Germany, using iron catalysts
at high pressure for oxygenate production, and cobalt catalysts at intermediate pressure for
liquid hydrocarbons and parafﬁn production. In general, the mechanism is a reductive
oligomerization of CO following a geometric progression in chain growth (Haggin, 1986).
The deﬁning reactions can be described by the formation of linear alkanes (and/or
alkenes),

nCO + 2nH2 ---> (012),, + nH2O

nCO + nH20---> nH2 + nC02

which sum to give the overall reaction,

2nCO + nH2 --> (CH2)n + nC02
for which the paraffin to oleﬁn ratio is mainly a function of the catalyst used and of the
initial CO/I-I2 ratio. One of the main limitations of conventional Fischer-Tropsch synthesis
technology is the lack of catalyst selectivity, with product fractions typically including a
C3 to C4 fraction, a gasoline fraction in the C3 to C10 range, and a diesel oil fraction as

27

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28

well (Haggin, 1981a). The wide product distribution is a functionof the type of catalyst
metal, which are normally Type VIII heavy metals, the pressure and temperature of
reaction, the CO/Hz ratio, and the presence of catalytic promoters (King et al., 1981).
Only the C1 compounds can be produced with 100% selectivity, i.e., methanol and
methane formation (King et al., 1981), although not until recently have detailed catalytic
studies elucidated the speciﬁc reaction mechanism for methanol synthesis (Chinchen et al.,
1990).

In general, the Fischer-Tropsch reaction can be compared to a chain polymerization
process and will thus exhibit a product distribution identical to a Schulz-Flory distribution
(King et al., 1981). Revitalization of Fischer-Tropsch technology has therefore focused on
the selectivity issue and has been based on catalyst improvement and development. These
efforts to more accurately control product distribution center on catalyst modiﬁcation, such
as the addition of zeolites to heavy metal catalysts (Haggin, 1981b). Another area is in the
use of catalyst promoters, either metal oxides like alumina or energetic promoters such as
alkali carbonates (Haggin, 1981b). Interception of surface-level intermediates during
reaction and physical regulation of hydrocarbon chain growth by catalyst pore size
restriction mechanisms are other exploratory efforts (King et al., 1981).

The only commercially proven Fischer-Tropsch synthesis plants for fuel production
using synthesis gas from coal are in South Africa, known as SASOL I, II, and 111. These
plants are geared predominantly towards production of gasoline and diesel fuels, but a
signiﬁcant amount of chemicals are also produced in both the gasiﬁcation and synthesis
steps (Dry, 1987). These plants use Lurgi gasiﬁers for a synthesis gas with a low CO/H2
ratio, use ﬂuidized bed reactors with iron catalysts for the Fischer-Tropsch synthesis,
remove sulfur and other contaminants using Rectisol technology, and employ nickel bed
catalysts in the steam reformation reactors (Dry, 1987). A generalized Sasol process
scheme, from Dry, (1987), is shown in Figure 4. Several speciﬁc factors attributed to the
success of the SAS OL plants include the availability of nearby, inexpensive coal, the lack

29

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

 

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30

of a domestic petroleum industry, and an industrial market far from the coast and thus far
from potential petroleum import stations (Dry, 1987). The speciﬁc economic and industrial
situation in South Africa is therefore the primary reason for SASOL's continued success.

One factor of concern in direct synthesis routes is the required CO/Hz ratio and the
corrresponding losses of CO2 and H20 during many synthesis reactions. Some

representative direct routes which are in various stages of development are shown in Table

2, along with the reaction stoichiometry and the required CO/H2 ratio. The commercial
potential of these direct synthesis routes, except for the methanol route, which is has been

proven commercially, is still being explored. Japanese researchers, as part of the Japanese

C1 Chemistry National Project, have done considerable development in both

catalysis/reaction control and reaction/process design for production of ethylene glycol,

Table 2. CO/Hz ratios of direct synthesis routes for chemicals from synthesis gas.

 

 

Product Reaction Stoichiometry CO/Hz Ratio
Methanol co + 2H2 ---> CH30H 1:1
Acetic Acid 2C0 + 2H2 ---> CH3COOH 1:1
Ethanol 2C0 + 4H2 ---> C2H50H + H2O 1:2
Propanol 3C0 + 6H2 ---> C3H7OH + H20 1:2
Ethylene 2C0 + 4H2 ---> CH2CI-12 + 2H20 1:2
Ethylene Glycol 2C0 + 3H2 ---> HOC2H40H 2:3

 

ethanol, acetic acid, and oleﬁns (Nakamura, 1990). Production of C1 to C6 alcohols is the
primary mission of the Research Association for Petroleum Alternative Development
(RAPAD), which is an international organization with a strong membership base in both
France and Japan. Up to 99% of C1 to C6 alcohols has been achieved in bench scale
tests, with up to a 30% weight fraction of C2 to C6 alcohols (Haggin, 1986). The

31

question of whether these syntheses are better performed in a direct or indirect fashion is
an ongoing issue, although the current technology greatly favors the indirect synthesis
routes; direct synthesis of these chemicals has been described as technology for the next
century (Ehrler and Juran, 1983).

A direct synthesis route which was originally a spin off of the Fischer-Tropsch
synthesis reactions are oxonation or hydroformylation reactions. These "oxo" routes are
also known as Roelen chemistry, after Otto Roelen, who invented the technology in
Germany over 50 years ago. The principle is based on reacting CO and H2 with an
unsaturated hydrocarbon, primarily olefms, to produce aldehydes and ultimately longer-
chain alcohols. Cobalt is generally the catalyst used, although iron, rhodium, and
ruthenium have exhibited similar properties (Haggin, 1981a). The aldehydes produced by
hydroformylation are most often hydrogenated to form alcohols, such as in the formation
of butyraldehyde and ﬁnally n-butanol and isobutanol from propylene; 2-ethylhexyl
alcohol is also produced by this route (Ehrler and J uran, 1983).

Another indirect synthesis route is carbonylation, in which CO reacts with unsaturated
and nucleophilic compounds to yield carboxylic acid derivatives. The typical catalysts for
these reactions are nickel, iron, and cobalt, and also rhodium, ruthenium, and palladium to
a lesser extent (Haggin, 1981a). The Monsanto process for acetic acid production from
methanol is an example of carbonylation. Besides methanol, other common nucleophiles
used in carbonylation are water, ammonia, alcohols, and carboxylic acids (Haggin,
1981a)

Many experts agree that indirect synthesis routes via methanol from synthesis gas
comprise the major future of using coal-derived gases for synthetic fuels and chemicals
production (Frank, 1982). The potential of methanol chemistry lies in both synthetic fuels
and chemicals production, including transportation fuels, diesel fuels, ethylene glycol,
acetic acid, acetic anhydride, and a host of others. The direct applications of methanol

itself are also of importance, both in its solvent and fuel properties.

32

One of the most recent triumphs in C1 chemistry development has been the Mobil
process for gasoline production from methanol, via zeolites. In the Mobil process, coal is
gasiﬁed to produce synthesis gas, which is then reacted to produce methanol at near 100%
selectivity by conventional methods. The methanol is then catalytically reacted to form
high octane gasoline and light hydrocarbons (C2 to C4 oleﬁns and C6 to C10 aromatics)
using the Mobil ZSM-S shape-selective zeolite catalyst (King et al., 1981; Keim, 1987).
This is a fully developed commercial technology, with operating plants in New Zealand
(Haggin, 1986). An extension of this technology is a two-stage approach for conversion
of methanol to olefms, and then selective conversion of the olefms to gasoline, a route
which produces more premium transportation fuels (Haggin, 1986). The methanol to
oleﬁn technology is of itself industrially important; under certain reaction conditions
ethylene and propylene can comprise almost the entire product spectrum (Haggin, 1981a).

For the the indirect synthesis of chemicals from synthesis gas via methanol, there are
basically 3 major routes: carbonylation, reductive carbonylation, and oxidative
carbonylation. In general, the normal carbonylation route is a commercial technology,
while the reductive and oxidative carbonylation routes are still under development (Keim,
1987). The Monsanto process for acetic acid production, as mentioned previously, is the
prime example of commercial technology for carbonylation and indirect methanol
chemistry. The reaction scheme is,

C0 + 2H2 --> CH30H
CH30H + CO ---> CI-I3COOH
and both reactions are highly selective and proceed under low pressure conditions (Ehrler
and Juran, 1983).

By reacting methanol and acetic acid, methyl acetate can be obtained, which upon
further reaction with CO will yield acetic anhydride. This indirect route is the basis for
commercial technology employed by the Tennessee Eastman Company, in Kingsport,

Tennessee. The process feeds a portion of the total coal-derived synthesis gas into a

33

cryogenic separater, where a pure CO stream is created for the acetic anhydride plant and
an H2-rich stream is created for adjustment of the CO/Hz synthesis gas ratio for methanol
production. The methanol is sent to a methyl acetate plant, where it is reacted with acetate
to form methyl acetate, which is then passed to the acetic anhydride plant to combine with
CO and produce acetic anhydride. The acetic anhydride is susequently reacted with
cellulose in an acetylation step to form cellulose acetate. This reaction releases acetate,
which is recycled to the methyl acetate plant. (Mayﬁeld and Agreda, 1986). The entire

process reaction scheme is basically,

Coal + 2H20 ---> C0 + H2 (1)
C0 + 2H2 ---> CH30H (2)
CH3COOH + CH30H ---> CH3COOCH3 (3)
CI-I3COOCH3 + CO ---> (CH3CO)2O (4)
(CH3CO)2O + Cellulose ---> Products + CI-I3COOH (5)

where reaction (1) is the coal gasiﬁcation step, reaction (2) is the methanol synthesis step,
reaction (3) is the methyl acetate synthesis step, reaction (4) is the acetic anhydride
synthesis step, and reaction (5) is the ﬁnal product formation step (Agreda, 1988). This
innovative chemistry is the result of 20 years of effort, and is a prime example of
technology diversiﬁcation from a process that was formerly based on natural gas and
petroleum fuel gases (Agreda, 1988).

There are several promising but commercially unproven technologies based on
reductive and oxidative carbonylation of methanol. Reductive carbonylation, or
homologation, can yield ethanol or acetaldehyde from methanol and CO/Hz, by the
reactions,

CH30H + C0 + 2H2 ---> C2H50H + H20

CH30H + C0 + H2 ---> CH3CHO + H20
which have yielded 80% selectivity for acetaldehyde over complex cobalt catalysts (Keim,
1987). Reductive carbonylation of methyl acetate can yield ethyl acetate or ethylidene

34

diacetate, which is a precursor to vinyl acetate (Haggin, 1986). Oxidative carbonylation of
methanol can yield dimethyl carbonate or dimethyl oxalate, which is a precursor to
ethylene glycol (Keim, 1987).

Other indirect routes include formaldehyde and methyl formate chemistry. Methyl
formate can be obtained by carbonylation of methanol, dehydration of methanol, methanol
oxydehydrogenation, by formaldehyde disproportionation, or by direct CO
hydrogenation. An application for methyl formate is as an intermediate in the synthesis of
formaldehyde, but other reactions are possible to yield a variety of ﬁne chemicals (Keim,
1987). Formaldehyde derived from synthesis gas can also be used as an intermediate in
indirect processing, with synthesis of compounds such as glycol anhydride, glycolic acid,
and methyl methacrylate possible. The potential, however, for formaldehyde as an
important intermediate in synthesis gas utilization is doubtful (Haggin, 1986; Keim,
1987).

11591°° ESl°E-D°1C I

The previous section has detailed some of the potential routes for synthesis gas in the
production of gaseous fuels (SNG), liquid fuels, electricity (IGCC), and a wide variety of
chemicals and chemical feedstocks obtained both by direct and indirect processing routes.
The ultimate application of these products and their processes will depend on their market
value, market share, and many other economic considerations. For the purposes of this
review, it sufﬁces to state that all of the listed commercial and near-commercial
technologies have existing markets and relatively clear applications as feedstocks,
intermediates, or ﬁnal products. The applications which are not entirely clear are for the
oxygenate fuels and their derivatives. It is these compounds that are the focus of this
section.

A major disparity in the U.S. energy situation as far as the extent of its diversiﬁcation is
in the transportaion sector. With 20% of the nation's energy consumption occuring in

cars, trucks, and buses and 98% of that energy being derived from petroleum, the

35

transportation industry is particularly sensitive to ﬂuctuations and shortages in the supply
and ﬂow of imported oil (Moore and Shadis, 1984). There has thus been ongoing research
and development pertaining to synthetic fuels and fuel additives, and their potential
application in the transportation sector to decrease the present petroleum-dependent nature
of the industry. But complete substitution of petroleum-based gasoline and diesel fuels
with synthetic fuels and alcohols will be a long-term endeavor. A more immediate problem
is the effort to increase octane ratings because of the phasing out of both lead and tetraethyl
or tetramethyl lead as octane enhancers.

The octane rating of a motor fuel is expressed as the octane number, a reﬂection of the
gasoline quality which most current automobile engines require in high amounts for
efﬁcient operation. Speciﬁcally, it is a reﬂection of the anti-knock property of a fuel,
expressed as a percentage of isooctane that must be mixed with n-heptane to equal the
knock intensity of the tested fuel. There are several kinds of octane numbers, including the
research octane number (RON), the motor octane number (MON), and the road octane
number (RdON). The RON is a test for fuels used in spark-ignition engines at low speed
and low temperatures. The MON is a test for fuels used in high speed, high temperature
conditions. The RdON is for intermediate applications, and is tested under actual driving
conditions. The average of the RON and MON values is the anti-knock index, or blending
octane number, and is used as a substitute value for an actual road test octane number. The
typical gas station pump octane value is the blending octane number.

The phasing-out of environmentally hazardous octane enhancers has created new
markets and potential applications for oxygenates such as alcohols and their ether
derivatives as clean burning blending agents, octane enhancers, and direct motor fuels.
These new, potentially coal-derived oxygenates include methanol, ethanol, n-butanol, and
ethers such as methyl tert-butyl ether (MTBE) and ethyl tert-butyl ether (ETBE). The
chemical and fuel properties of these compounds compared to conventional gasoline motor

fuels are shown in Table 3. One of the most favorable properties of oxygenates is their

36

Table 3. Properties of various fuel oxygenatesa.

 

 

Product Blending AHcomb Boiling Point
RON MON Octane (kBtu/lb) (0C)
Methanol 111 91 101 9.77 64.6
Ethanol 108 91 100 11.59 78.5
n-Butanol 101 87 94 15.02 1 17
MTBE 116 99 108 17.22 55.4
ETBE 117 105 111 18.79 72.8
Gasoline 91 83 87 20.13 Range

 

aFrom Ecklund and Mills, 1989a; Iborra et al., 1988.

octane number. The other more favorable properties of alcohols and ethers are their high
miscibility in gasoline and higher blending octane values compared to conventional motor
fuels.

The application of oxgenates as octane enhancers depends on several factors, including
the miscibility in gasoline, the vapor pressure, the blending octane number, and the burn
characteristics. When comparing various oxygenates, methanol is generally the base of
comparison. The ethers, which are formed from the reaction of either methanol or ethanol
with isobutylene, have greater octane values and higher heats of combustion compared to
methanol (Piel, 1988; Ecklund and Mills, 1989b; Iborra, 1988). The higher alcohols, in
the C2 to C4 range, also have higher heats of combustion, higher octane values, and can
also inhibit phase separation when used as cosolvents in methanol-gasoline mixtures
(Ecklund and Mills, 1989b). In general, there is great potential for increased usage of
oxygenates as gasoline octane enhancers, extenders, and blending components for use

with motor fuels.

37

W
The development of coal gasiﬁcation technology, of integrated coal-based power and

chemical plants, and of synthesis gas reactor and catalyst technology has served to
revitalize and refocus the synthetic fuels and chemicals industry, even in the advent of
inexpensive petroleum. The cultural bias against coal and coal-ﬁred plants has to a certain
extent been steadily eroded due to adherence to strict environmental compliance issues,
long-term uncertainty and instability of petroleum supplies, and the potential economic
advantages arising from a diversiﬁed industry based on domestic resources and labor.

However, the obstacles to commercialization are somewhat more complex than mere
social, political, and environmental issues. Long-term dependence on petroleum has
created an infrastructure committed to that resource, and which will require signiﬁcant
restructuring and far-sighted investment in order to diversify the energy supply base in the
face of short-term, proﬁt-directed corporations, both on the supply (petroleum companies)
and the demand (automotive and transportation companies) sides of the industry. Without
the assistance of another major crisis in the ﬂow of petroleum imports, these obstacles will
be difﬁcult to surmount. The potential for diversiﬁcation is not in doubt, as a look at the
global examples in South Africa (commercial Fischer-Tropsch synthesis from coal) and in
Brazil (ethanol from biomass) will attest. Yet the factors arising to the success of these few
and often cited examples were not only technical, indeed, the supplies of coal in South
Africa and biomass in Brazil were a large piece of the success equation, as were
government subsidies and public acceptance. It therefore can be concluded that for
successful commercialization of coal-based processes, the technical, political,
environmental, and social climates must all be condusive.

Methanol from coal, although by no means a synthetic fuels and chemicals panacea, has
been one of the major developments in recent times. Finding new uses for methanol and
methanol derived products, as well as serving existing ones, has been an important aspect

of the technology development. The current outlook for methanol is considered to be very

38

promising (Frank, 1982; Ecklund and Mills, 1989a). One aspect of the commercial
promise of methanol as a product or as an intermediate is the speciﬁcity of its formation
from synthesis gas, with almost 100% yields obtainable, as described previously. This
speciﬁcity is also apparent in other areas of C1 chemistry, and will continue to be a major
advantage of indirect synthesis gas processing routes via methanol (Ehrler and Juran,
1983).

In order to substitute an existing petroleum based product or industry with a coal-based
product or industry in the absence of a cohesive and focused national energy strategy, the
potential niches for application will have to be well worked out, as can be seen in the
examples at Tennessee Eastman and the Great Plains Coal Gasiﬁcation project.
Alternatives to existing fuels and chemicals that are presented as environmentally sound
will be received favorably in the existing industrial climate, as are the above examples. It is
thus in this light that clean-burning, alternative motor fuels and fuel additives have future
potential, ﬁrst by diversifying the energy sources for such products (coal and biomass
versus petroleum), and second, by presenting environmentally "friendly" alternatives

(oxygenates versus lead compounds) to the industry.

1.2. Microbial Metabolism of Synthesis Gas Components

1212. .1 [1-21111].

The ongoing discovery of l-carbon metabolizing bacteria and their diversity in nature
has led to a somewhat confusing and sometimes redundant terminology for classiﬁcation
of the organisms. Autotrophy refers to organisms which use CO2 as their sole carbon
source and which grow in the absence of organic matter. Heterotrophy refers to organisms
which grow on organic carbon and energy sources. Methylotrophy refers to organisms
capable of growth on organic compounds containing no carbon-carbon bonds, such as
methanol. Mixotrophy refers to cosubstrate metabolism. When classifying microrganisms

by their carbon sources, autotrophs and heterotrophs are presently the accepted terms, and

39

thus classiﬁcation of CO-utilizing organisms requires, either further deﬁnition or a
broadening of the current classiﬁcations. The term unicarbonotrophy has been suggested
for organisms capable of growth on l-carbon compounds as their sole carbon and energy
sources (Zeikus, 19833).

For the purposes of this review, the term unicarbonotroph will be used for organisms
capable of growth on CO, and the use of the terms autotroph and methylotroph will be
restricted to their literal deﬁnitions above. This terminology emphasizes the great diversity
of several bacterial species, outlined below, such as the anaerobic bacterium
Butyribacten'um methylotrophicum, which is capable of heterotrophic growth on glucose,
autotrophic growth on H2 and CO2, methylotrophic growth on methanol,
unicarbonotrophic growth on C0, and mixotrophic growth on CH30H and C02.

The major mechanisms for CO2 ﬁxation in autotrophic microorganisms include the
ribulose diphosphate carboxylase cycle, or Calvin cycle, the reverse di- or tricarboxylic
acid cycles, or reductive carboxylic acid cycles, and the acetyl-CoA pathway, which is a
uni-directional mechanism. The Calvin cycle involves ﬁxation of CO2 via ribulose
diphosphate, yielding 2 molecules of 3-phosphoglyceric acid for every CO2 molecule
consumed. The 3-phosphoglyceric acid is then further reacted in an ATP and electron-
consuming reaction to form glyceraldehyde-3-phosphate, as well as to regenerate ribulose
diphopshate, which is the intermediate used for subsequent catabolic reactions, including
cell synthesis (Brock et al., 1984). This cycle, a C3 cycle, functions in all aerobic
photosynthetic organisms, in some anaerobic phototrophs, and in several
carboxydotrophs, or aerobic CO-utilizing bacteria (Meyer and Schlegel, 1983).

The reverse di or tricarboxylic acid cycles are C4 pathways which use acetyl-CoA as
the initial CO2 acceptor through ferrodoxin-mediated carboxylation to pyruvate, and then
further phosphorylation and carboxylation to oxaloacetate. In the reverse dicarboxylic acid
cycle, the oxaloacetate is then further reduced, either to malate, fumarate, or succinate, and

the acetyl-CoA regenerated to complete the cycle. The net CO2 ﬁxation is thus 2 moles per

40

mole of oxaloacetate formed. In the reverse tricarboxylic acid (TCA) cycle, the carbon and
electron ﬂow proceeds with more carboxylation reactions from oxaloacetate ultimately to
citrate, which is cleaved to regenerate acetyl-CoA and oxaloacetate, resulting in a net C02
ﬁxation of 4 moles per mole of oxaloacetate formed (Brock et al., 1984).

The acetyl-CoA pathway has been studied in both acetogenic and methanogenic
bacteria, and is the fundamental metabolic pathway operating in the bacterium used in the
present study. This mechanism is basically a catabolic reduction process, where reducing
equivalents from a variety of dehydrogenated compounds (Hz, H20) are used to form the
methyl and carboxyl groups of acetate, via acetyl-CoA as an intermediate (Fuchs, 1986).
The ubiquitous enzyme in this transformation is carbon monoxide dehydrogenase
(CODH), also known as acetyl-CoA synthase, a multi-component enzyme complex of
varying homlogy in several species of acetogenic bacteria (Fuchs, 1986). Although
initially conceived by Wood over 40 years ago, the elucidation of this pathway for both
CO and CO2-utilizing bacteria has been ongoing, particularly in the last decade, and
current research trends are focused on the energetics of acetogenic metabolism and the
mechanism for ATP generation during autotrophic and unicarbonotrophic growth.

A variety of microorganisms representing a diverse cross section of the bacteria
division of the kingdom procaryotae are capable of l-carbon compound metabolism,
primarily by one of the three major mechanisms of carbon ﬁxation listed above. However,
in the context of potential industrial application for conversion of the major synthesis gas
components, the characteristics of anaerobic bacteria, specifically unicarbonotrophic
anaerobic bacteria, represent the most promising route for production of fuels and
chemicals (Datta and Zeikus, 1982; Zeikus, 1983b; Linden, 1988). Anaerobic metabolism
conserves the chemical energy content of the substrate very efﬁciently in the abscence of
O2 and oxidation, a characteristic which provides unicarbonotrophs with a metabolic
efﬁciency greater than all other biosynthetic life forms (Zeikus, 1983b). The autotrophic

and unicarbonotrophic organisms outlined in the following pages thus represent

41

speciﬁcally those organisms capable of anaerobic oxidation of CO or anaerobic utilization
of H2 and C02. In the majority of these organisms, particularly the acidogenic bacteria, it
is the Wood or acetyl-CoA pathway that is the mechanism for l-carbon compound
utilization, and thus this pathway is discussed in more detail below.
guru an..." guru but an - am": new M; :r. 'o:

Attaining a unifying theme for organisms capable of metabolizing the primary synthesis
gas components CO, H2, and CO2 is a difﬁcult problem because of the phylogenetic
diversity of the many species. For this review, the anaerobic bacteria capable of such
metabolism are classiﬁed as either methanogenic, sulphidogenic, phototrophic, or
acidogenic. The emphasis is thus on metabolic similarity as opposed to taxonomic or
phylogenetic similarity.

Methanogenic bacteria are obligate anaerobes and members of the procaryotic group
Archaebacteria, family Methanobacteriaceae. Typically, C02 is the terminal electron
acceptor and H2 the electron donor during methanogenesis. These organisms are
considered to play an important role in the carbon cycle, with approximately 80% of the
atmospheric methane believed to have a biological origin (Brock et al., 1984). There is no
direct evidence for the existence in methanogens for either the Calvin cycle or
cytochromes, and novel metabolic mechanisms have been proposed which include a
propensity of corrinoid species, unique coenzymes, and coenzyme factors (Zeikus, 1983b;
Zeikus et al., 1985; Linden, 1988). Several methanogenic species of bacteria are capable
of metabolizing CO and/or H2 and CO2, abilities which have in most cases been achieved
only after long periods of culture adaptation (Zeikus, 1983b). In general, the energy-
yielding metabolism of methanogenic bacteria is deﬁned by the oxidation of H2 and the
reduction of C02 to form methane. The most well characterized species capable of
unicarbonotrophic growth are in the genera Methanobacterium and Methanosarcina, which
are genera readily distinguished by cell morphology and ultrastructure analysis (Zeikus
and Bowen, 1975).

42

Methanobacterium thermoaurotrophicum and Methanobacterium formicicum are rod-
shaped, gram-positive bacteria with a high guanine and cytosine (G+C) content in the
DNA (>50%) and numerous intracytoplasmic membranes (Zeikus, 1977). Both of these
organsims are nonmotile and use H2 and CO2 as substrates for methanogenesis.
Methanobacterium thermoautotrophicum, a thermophilic, ﬁlament-forming organism
isolated from heat exchanger sewage sludge (Zeikus and Wolfe, 1972), is considered to be
the most proliﬁc methogen known, and but is not an obligate autotroph, since growth on
acetate has been observed (Linden, 1988). Doubling times as low as 2 hr have been
observed on H2 and C02 in ammonium-limited continuous culture studies (Kenealy et al.,
1982). Adaptation of M. themoautotrophicum to grow with CO as the sole carbon source,
with following the reaction

4C0 + 2H20 --> CH4 + 3C02
has been achieved in CO atmospheres up to 30%, albeit at substantially reduce rates of
growth (approximately 1%) compared to H2 and CO2 (Daniels et al., 1977). Growth was
completely inhibited at C0 partial pressures approaching 60% during these studies
(Daniels et al., 1977).

Another species of bacteria, which has been described as the most metabolically
versatile methanogen known, is Methanosarcina barkeri. This bacterium is a gram-
positive, packet—forming cocci, with a G+C content over 50% (Zeikus, 1977). M. barken',
of which there are several distinct strains, is capable of either autotrophic,
unicarbonotrophic, or mixotrophic growth on a variety of compounds as either the sole
carbon or sole carbon and energy souces. Autotrophic growth on H2 and CO2 or
methylotrophic growth on CH30H has been observed with doubling times of 11 hr and 8
hr, respectively (Weimer and Zeikus, 1978). Mixotrophic growth on H2/CO2/CH3OH is
also possible during methanogenesis (Weimer and Zeikus, 1978). Methane formation
from acetate has also been observed, but with almost a 40 hr doubling time and an

extremely long adaptation period (Zeikus, 1983a). Growth on C0 as the sole carbon

43

source with a doubling time of 65 hr in the presence of H2 has been reported (O'Brien et
al., 1984).

1-Carbon compound metabolism in sulphidogenic bacetria, which comprise an
extremely diverse group of organisms, has been documented speciﬁcally for CO or H2
and CO2, although the speciﬁc physiological concepts have not been well determined.
Such organisms include Desulfovibrio vulgaris, Thermodesulfotobacterium commune,
and Desulfovibrio desulfuricans (Zeikus, 1983a). Growth of D. vulgaris with CO as the
sole electron some for cell mass production and sulfate reduction was achieved at partial
pressures up to 4%, but was inhibited at greater concentrations of CO in the gas headspace
(Lupton et al., 1984). Uptake of dilute CO gas has also been observed by D.
desulﬁtricans, but at very low rates of metabolism and with no apparent energy-generating
mechanism (Uffen, 1981). Thennodesulfotobacterium commune has exhibited growth on
H2 and CO2 in the presence of acetate, although prolonged periods of cultural adaptation
were required (Zeikus, 1983a). In general, metabolism of synthesis gas components in
sulphidogenic organisms, although having been documented in a few speciﬁc cases, is not
nearly as deﬁned as in methanogenic or acetogenic bacteria, and is only included here in
order to exemplify the diversity of species capable of such metabolism.

Similarly, photosynthetic bacteria such as Rhodospirillum rubrum and
Rhodopseudomonas gelatinosa have also been documented with CO metabolizing ability
following the overall water-gas shift reaction,

C0 + H2O ---> H2 4» CO2.
R. rubrum is known to metabolize CO gas anaerobically and in the dark, although growth
cannot be supported on CO alone (U ffen, 1981). R. gelatinosa , also under dark
conditions, is capable of growth with CO as the sole carbon and energy source, although
addition of trypticase was required to achieve the reported doubling time of 6.7 hr (U ffen,
197 6). As for the sulphidogenic bacteria described above, CO metabolism in phototrophs
is also poorly characterized and not well understood.

44

The occurence of l-carbon metabolism is perhaps best-characterized and understood in
acidogenic, anaerobic bacteria, primarily acetogenic bacteria in the family Bacillaceae
genus Closm'dium and family Propionibacteriaceae genus Eubacterium. Unicarbonotrophy
has only recently been recognized in homoacetogens, and in general, acidogenic bacteria
are the most interesting because of the potential for acetate, ethanol, butyrate, and butanol
synthesis from C0 or H2 and CO2 found in this diverse collection of microorganisms. It
is important to stress the distinction between the terms acidogen and acetogen, which are
often confused and thus misused. Acidogen refers to the products from metabolism, in
this case mainly acetic acid, with a usage similar to that of the term methanogen. Acetogen
is a term which speciﬁcally refers to catabolic pathways which are based on formation of
acetyl-CoA as an intermediate, with homoacetogens being organisms with acetyl-CoA as
the sole intermediate for all catabolic reactions, whether from heterotrophic, autotrophic,
or unicarbonotrophic mechanisms.

Closm'dium thermoacezi cum, a spore-forming thermophile, is probably the most
intensely studied and well known homoacetogenic bacterium. This organism has a G+C
content of 54%, and grows on both CO, with a 16 hr doubling time, or H2 and C02, with
an 18 hr doubling time, but is not capable of growth on CH30H (Kerby and Zeikus,
1983). Growth on CO was reported not to be inherent but was obtained after a prolonged
adaptation period (Kerby and Zeikus, 1983). Interestingly, the capacity for CO oxidation
was ﬁrst observed during heterotrophic growth on glucose (Diekert and Thauer, 1978).
Both growth on C0 and on H2 and CO2 yielded acetic acid, with theoretical product
formation stoichiometries of

4C0 + 2H20 ---> CH3COOH + 2CO2
4H2 + 2CO2 --> CH3COOH
calculated from batch growth results (Kerby and Zeikus, 1983). Several key enzymes and
electron carriers have been discovered and characterized in C. thermoaceticum, including

the ﬁrst evidence for heme-prosthetic groups, such as cytochromes and menaquinones, as

45

opposed to strictly corrinoid-based electron carriers (Gottwald et al., 1975), a discovery
which still has widespread implications when discussing electron ﬂow in acetogenic
organisms. Recent evidence suggests that the menaquinone in C. thermoaceticum is the
same as for Clostridr’um thermoautotrophicum, discussed below (Das et al., 1989).
Electron carrying ferrodoxin proteins have also been isolated (Elliott and Ljungdahl,
1982), as has a nickel—based CO dehydrogenase enzyme (Drake et al., 1980) and
hydrogenase enzyme (Drake, 1982b).

Several more recent physiological discoveries with C. thermoaceticum include the
existence of a membrane bound ATPase complex, supporting the increasingly widespread
belief that homoacetogenic bacteria require active electron-transport phosphorylation
mechansisms. This ATPase complex, which functions based on the proton-motive force,
has been characterized and puriﬁed to homogeneity (Mayer et al., 1986; Ivey and
Ljungdahl, 1986). Membrane vesicles containing CO dehydrogenase have also been
shown to generate a proton gradient when exposed to CO, which is another piece of
physiological evidence for a complete electron-transport phosphorylation mechanism in C.
themroaceticum (Hugenholtz and Ljungdahl, 1989).

Other Clostridia have also shown either CO or H2 and CO2 metabolizing ability.
Clostridium pasteurianum exhibited CO oxidation in cell-free extracts (Thauer et al.,
1974), and while growing on glucose (Fuchs et al., 1974), although growth with CO as
the sole carbon source was not reported. The CO dehydrogenase enzyme was determined
to be nickel-dependent in C. pasteurianum, and similar to that present in C.
thermoaceticum (Diekert et al., 1979; Drake, 1982a). Closm’dium thennoautotrophicum
has a G+C content of 54%, and exhibits an autotrophic metabolism similar to that of C.
thermoaceticum, with growth on H2 and CO2 observed at a doubling time of 8 hr and at
the same 4:2:1 H2:CO2:acetate stoichiometry (Wiegel et al., 1981). Growth on CH30H
was also reported, as was a G+C content of approximately 54% (Wiegel et al, 1981).
Similarly, CO dehydrogenase and hydrogenase enzyme were found to play important roles

46

in acetate synthesis from H2 and C02 (Clark et al., 1982). Another species exhibiting H2
and CO2 metabolism is Clostridium aceticum, which has a G+C content of 33%, and but
does not exhibit growth on CH30H (Braun et al., 1981). A recently discovered bacteria,
now known as Closm'dium ljungdahlii, was originally isolated from chicken waste, and is
capable of both ethanol and acetate synthesis from C0 gas (Barik et al., 1988). However,
the speciﬁc physiological characteristics of this organism have not yet been determined.

Acetobacterium woodii is a rod-shaped, mesophilic bacterium isolated from a marine
environment, which possesses a G+C content of 39%, and is capable of growth on CO,
H2 and CO2, and CH30H and CO2 to produce acetate (Tanner et al., 1978; Kerby et al.,
1983). The l-carbon metabolism in A. woodii is considered to be very similar to that of C.
thermoaceti cum because of the presence of corrinoids, a nickel dependent CO
dehydrogenase, and other key enzymes associated with l-carbon transformations in
acetogens (Ragsdale et al., 1983b; Kerby et al., 1983).

Peptosrreptococcus productus strain U-l was originally isolated from sewage digestor
sludge and is perhaps the fastest growing bacteria exhibiting CO metabolism, with a 1.5 hr
doubling time in a 50% CO atmosphere and doubling times less than 2 hr for CO
atmospheres approaching 90% (Horowitz and Bryant, 1984). This organism produces
acetate and CO2 from CO, and is also capable of H2 and CO2 metabolism to form acetate
with a doubling time of 5 hr in an 80% atmosphere (Horowitz and Bryant, 1984). Growth
on methanol has not been observed, nor has growth on formate (Ljungdahl, 1986). P.
productus is one of the few CO-utilizing baceteria for which extended batch and
continuous fermentations have been conducted, including extensive fermentation kinetic
analyses during acetate production from C0 in both batch and chemostat sytems (Vega et

al., 1989a; Vega et al., 1989b).

A sheep rumen isolate, Eubacterium Iimosum is capable of both acetate and butyrate

production depending on the l-carbon substrate. This organism is considered to be a

newer isolate of what was formerly known as Butyribacterium rettgeri. When grown on

47

H2 and CO2, E. limosum exhibits a 14 hr doubling time, producing acetate and minor
amounts of butyrate (Sharak Genthner et al., 1981). Faster growth (7 hr doubling time)
was observed on CH30H with equimolar production of acetate and butyrate, but only in
the presence of acetate and CO2 (Sharak Genthner et al., 1981). The ratio of CH30H to
CO2 is critical in determining the product ratio, regulation occuring speciﬁcally due to the
bicarbonate ion concentration (Pacaud et al., 1985). Growth on 50% CO was observed
with a 7 hr doubling time and sole production of acetate, this doubling time increased to 18
hr at a CO concentration of 75%, which was the highest partial pressure tested (Sharak
Genthner and Bryant, 1982). Butyric acid fermentation schemes using CH30H and CO2
have been proposed using E. limosum, schemes which depend on feedback inhibition
mechanisms, optimization of substrate ratios, and ultraﬁltration technology for
homobutyric acid production (Loubiere et al., 1986).

Probably the most versatile of unicarbonotrophic, anaerobic bacteria, Butyribacrerium
methylotrophi cum is capable of both acetate and butyrate production depending on the
substrate used. Growth on H2 and CO2 produces acetate with a 9 hr doubling time, while
growth on CH30H in the presence of acetate and CO2, also at a 9 hr doubling time,
produces butyrate in a 4:1 methanolzbutyrate ratio (Zeikus et al., 1980; Lynd and Zeikus,
1983). Product formation and fermentation rates in B. methylotrophicum during growth
on CH30H and CO2 were also found to be dependent on the substrate ratio, as for E.
Iimosum (Datta and Ogletree, 1983). A mutant strain of this organism, the CO strain, was
adapted to grow with a 12 hr doubling time on 100% CO gas, producing acetate and minor
amounts of butyrate (Lynd et al., 1982). The speciﬁc origin, physiology and biochemistry
of B. methylotrophicum are discussed in more detail in the next section.

12313"! E! '1111'

The speciﬁc mechanism of acetogenic metabolism in both the unicarbonotrophic and

autotrophic pathways of C. thermoaceticum and B. methylotrophicum have recently been

elucidated, particularly the role played by CO dehydrogenase in the formation of acetyl-

48

CoA (Zeikus et al., 1985; Kerby and Zeikus, 1987b; Fuchs, 1986; Wood and Ljungdahl,
1991). Figure 5, adapted ﬁom Zeikus et al., (1985), and Wood and Ljungdahl, (1991), is
a condensed outline of the autotrophic acetyl-CoA pathway for CO and CO2 utilization by
these organisms. Key steps for CO metabolism include the reaction of CO and H20 with
CO dehydrogenase to yield C02, the production of formate from C02 via formate
dehydrogenase, several tetrahydrofolate mediated transformations ultimately resulting in a
methyl—corrinoid complex, the reaction of CO with CO dehydrogenase to form a bound
carbonyl moeity, and ﬁnally the synthesis of acetyl-CoA, from both the methyl and
carbonyl groups bound to the CO dehydrogenase complex. Metabolism of H2 and CO2
occurs by the same mechanism, with H2 being used to reduce C02 to bound CO, or
[CO], in a ferredoxin dependent reaction, and the resulting carbonyl group reacting with
the methyl-corrinoid species via CO dehydrogenase, to ultimately form acetyl-CoA.

The speciﬁc reactions and the mediating enzymes and enzyme complexes are detailed in
Table 4. It is important to note the versatility of the CO dehydrogenase complex in that it
potentially mediates 5 distinct reactions, including the oxidation of CO to CO2, the binding
of CO, the reduction of CO2 to form a bound [CO] group, the binding of a methyl group
from the corrinoid protein, and the condensation reaction of the carbonyl group, the
methyl group, and the CoA ester to form acetyl-CoA.

The properties of carbon monoxide dehydrogenase have been recently reviewed
(Diekert et al., 1985), and the enzyme has been puriﬁed from both C. thermoaceticum and
A. woodii (Ragsdale et al., 1983a; Ragsdale et al., 1983b). The 440,000 MW enzyme
isolated from C. thermoaceticum contains 2 different subunits of three c0pies each in the
form 3383, with molecular weights of 78,000 and 71,000, respectively (Ragsdale et al.,
1983b). The enzyme is a nickel-containing, iron-sulfur protein which reduces ferredoxin
and cytochrome b in the presence of CO (Drake et al., 1980; Drake, 1982a). The iron
exists in the active site as clusters of Fe2S4, which are reduced by reaction with CO and

oxidized by reaction with CO2 (Diekert et al., 1985). The condensation of bound

49

CO CO

1 2H+2H+ H20 [CO]

 

 

 

 

 

 

 

21'] ATP + H20
ADP + Pi
[HC OH]
411 \1
[CH5 OH]
\ HCoA J
H20
Acetyl-CoA
Pi
HCoA
* Acetyl-P
| Cell Mass |
ADP
ATP
Acetate

 

Figure 5. General pathway for unicarbonotrophic, acetogenic
metabolism.

50

Table 4. Reactions in the unicarbonotrophic and autotrophic pathways for acetyl-CoA
synthesis in the acetogens B. methylotrophicum and C. thermoaceticuma.

 

 

Reaction Mediating Enzyme
Formation of the Methyl Groupb:
C0 + H20 ---> CO2 + 2H‘*‘ + 2e' Carbon Monoxide Dehydrogenase
CO2 + 2e” ---> HCOOH Formate Dehydrogenase
HCOOH + ATP + THF --> HCO-THF + ADP + H2O Formyl-TT-IF Synthetase
HCO-THF + H+ ---> HC-THF Methenyl-THF Cyclohydrase
HC-THF + 2e' ---> H2C-TI-IF Methylene-THE Dehydrogenase
H2C-THF + Fdz2e' ---> H3C-THF Methylene-THF Reductase
H3C-THF + Co-E ---> H3C-Co-E Corrinoid Protein
Formation of the Carbonyl Group:
C0 + CODH ---> CO-CODH Carbon Monoxide Dehydrogenase
H2 + Fd ---> Fd:2e' + 2H+ Hydrogenase
Fd:2e' + CO2 + CODH --> CO-CODH + Fd Carbon Monoxide Dehydrogenase

Formation of the Acetyl—CoA Group:
H3C-Co-E + CO-CODH ---> H3C-CO-CODH + Co-E
H3C—CO-CODH + CoASH ---> CH3CO-CoA Carbon Monoxide Dehydrogenase

 

aDerived from Zeikus et al., 1985; Kerby and Zeikus, 1987b; Wood and Ljungdahl,
199 1 .

bTHF = tetrahydrofolate; Fd = ferredoxin; Co-E = Corrinoid protein; CODH = Carbon
monoxide dehydrogenase.

carbonyl, methyl, and CoA groups catalyzed by this enzyme has recently been veriﬁed in
several different bacteria, including C. thermoaceticum (Ragsdale and Wood, 1985),
Methanosarcina thermophila (Abbanat and Ferry, 1990), and A. woodii (Diekert and
Ritter, 1982).

The other major components in the acetogenic pathway are tetrahydrofolate and
corrinoid protein(s). Tetrahydrofolate participates in the reduction of formate to a methyl
group, by serving as the carrier during catalysis by the 4 sequential enzymatic reactions;
these reactions are reversible depending on the substrate utilized (CO/CO2 or CH30H)
(Zeikus et al., 1985). The corrinoid protein is a cobalt containing carrier protein, also

known as cobamides, with 3 reduction states, and participates not only in the methyl

5 1
group transfer to CO dehydrogenase, but also in amino acid synthesis and in ferredoxin

oxidation (Ljungdahl, 1986).

The principles of metabolic energy generation in acetogens are based on both substrate-
level (SLP) and electron-transport (ETP) phosphorylation mechanisms. Energy
conservation by SLP is a mechanism which couples exergonic chemical reactions to the
synthesis of energy-rich intermediates such as thioeters or acid anhydrides, the hydrolysis
of which allow formation of ATP from ADP and inorganic phosphate (Thauer et al.,
1977). In autotrophic, acetogenic bacteria, this energy-rich intermediate is primarily acetyl-
CoA. However, metabolism of CO or H2 and CO2 to produce acetate does not produce
any net ATP via SLP because an ATP is consumed in the activation of the
formyltetrahydrofolate synthetase reaction for every acetyl-CoA formed, and an ATP is
produced for every acetate formed from acetyl-CoA, as shown in Table 4. For production
of cell mass, another source of ATP generation must be present, and a chemiosmotic
mechanism (ETP) to couple the proton motive force to ATP synthesis has often been
hypothesized to account for the ATP generation required for cell growth and other
metabolic reactions (Zeikus et al., 1985; Fuchs, 1986). The basic mechanism of electron-
transport phosphorylation is given in Figure 6, where soluble, membrane bound, and
intermembrane electron carriers facilitate the ejection of protons outside of the cell into the
environment, generating a proton gradient across the cell membrane. This gradient, or
proton-motive force, can then serve as the driving force to allow protons to re-enter the
cell, the potential energy of which is captured in the form of ATP synthesis. The electron-
transport pathway thus serves to link catabolic oxidation and reduction reactions to the
proton-motive force. Both physiological and structural results supporting such a
mechanism in acetogenic bacteria have been increasing over the last 6 years, and an ATP
synthase and corresponding electron carriers have been isolated in C thennoaceticum, as
described previously. Soluble and membrane-bound electron carriers have been isolated

from B. methylotrophicum, including ferredoxins, rubredoxins, cytochrome b and

52

maid: mm Outside
Low [11"] High [n+1
Catabolic Metabolism

Oxidation-Reduction

 

Electron Flow --->

 

 

 

 

 

 

\
\
\
Oxidation-Reduction ‘ A H+
I ‘
<--- Electron Flow <— — —
+
v H

ATP

 

 

Figure 6. General mechanism for electron transport-coupled
phosphorylation (ETP).

53

menaquinones, but the complete electron-transport system has yet to be ascertained
(Zeikus et al., 1985; Saeki et al., 1989a; Saeki et al, 1989b).

Unfortunately, without the complete picture for energy generation in autotrophic and
unicarbonotrophic acetogens, the speciﬁc energy yields for various substrates and
metabolic reactions will be very difﬁcult to verify, as discussed in later chapters.
Furthermore, other mechanisms of ATP synthesis are possible, including transport-
coupled phosphorylation (TCP), which as of this writing has only one known example in
P. modestum (J. Kemner, pers. comm.). Such a hypothesis would inevitably raise
questions of the possibilities for other, non-proton based ATP synthase enzymes, such as
sodium based ATPases. These questions are currently under investigation; C.
thermoaceticum has been determined to be sodium-independent during growth, but A.
woodii has been found to be quite sodium-dependent during growth (Wood and
Ljungdahl, 1991). In summary, it is now generally accepted that ETP mechanisms do
exist in homoacetogens, and that they play an important role in energy generation, but the
speciﬁc mechanisms and electron ﬂow pathways have yet to be elucidated in many
bacteria, as do the membrane-associated ATP synthase components.

1 2 | E . l E l' . E E . E .

The traditional feedstocks for anaerobic fermentation have been carbohydrates, but a
resurgence in industrial coal utilization would provide ample and inexpensive l-carbon
feedstocks for fermentation by selected unicarbonotrophic and autotrophic organisms. The
major synthesis gas components, CO, H2 and CO2, and also CH30H, are all readily
available through coal gasiﬁcation, or coal gasiﬁcation and reaction in the case of CH30H.
A general schematic for potential uses of l-carbon fermentation technology in processing
these compounds is shown in Figure 7, with several routes available for production of
fuels, chemicals, and other feedstocks or commodity biological products.

The biological conversion of synthesis gas to methane, analogous to industrial catalytic

SNG production from coal via steam reformation, is a reaction well- suited to some of the

54

 

W

Coal
Petroleum
Natural Gas

 

 

 

 

W
H2/C02 , co, CH30H, CH4

 

 

 

 
     
   
   

1-CARBON
Anaerobic

Fermentation

Technology

 

mm mm chm
M Eula Vitamins

Solvents Alcohols Egg“
Acids/Gases SNG Proteins

 

 

 

 

 

 

 

 

 

Figure 7. Potential biological routes for fuels and chemicals production from synthesis
gas components.

55

methanogenic organisms listed previously. Such reactions are typically referred to as
biomethanation processes; inherent in this deﬁnition is a potential for secondary product
formation such as short-chain fatty acids. The basic tenet of biomethanation is upgrading
the Btu content of synthesis gas, and studies with mixed cultures have shown CO2
formation from CO, and CH4 formation from H2 and CO2 (Wise et al., 1978).
Enhancement of methane product yields, bacterial growth rates, and improvement of rate
limiting metabolic steps including gas-liquid mass transfer have been identiﬁed as areas
requiring more detailed study (Zeikus, 1983b), an indication that this technology is still in
the exploratory stages of development. Regardless of the potential for biotechnology in
this area, competition with available natural gas reserves and with industrially available
routes for SNG production from coal could seriously limit the application of
microorganisms for methane production from synthesis gas.

Acidogenic bacteria have the potential for production of short-chain organic acids, 2-
and 4-carbon alcohols, and perhaps even single-cell protein from synthesis gas
components. The potential for acetic or butyric acid production by acetogens has been
well-documented, and product yields are particularly favorable (Zeikus, 1980). However,
acetic acid is presently a large-volume industrial chemical, and commercial technology for
acetic acid production from C0 and CH30H is readily available, as discussed previously
for the Monsanto process. Butyric acid, although not produced in nearly as large quantities
as acetic acid, is a potential precursor to butanol, or for esteriﬁcation in order to produce
other chemical intermediates (Datta, 1981).

The recent discoveries of direct ethanol and butanol production from synthesis gas
components via unicarbonotrophic, anaerobic bacteria have created a potential application
for such organisms in the production of fuels, fuel additives, and industrial solvents.
Ethanol has current potential as a direct gasoline additive, in addition to its chemical uses
and solvent properties. Butanol is a so-called plasticizer alcohol, and can serve either as a

chemical feedstock for production of butyl acetate or butyl acrylate, or can be potentially

56

be used as a gasoline octane enhancer or fuel additive. In general, application of l-carbon
metabolizing anaerobic bacteria for conversion of synthesis gas could ultimately depend on
the discovery of a speciﬁc industrial niche based on the 4 most readily available products
acetate, butyrate, ethanol, and butanol, or more likely on processes using these products

as intermediate chemical feedstocks.

1.3. Physiology and Metabolism of Butyribacterium methylotrophicum

1310.. Hi 1C1 .. [1 [205°

The original strain of Buryribacterium methylotrophicum (ATCC33266), known as the
Marburg strain, was isolated by Zeikus and coworkers in 1980 from sewage digestor
sludge in Marburg, Federal Republic of Germany. B. methylotrophicum is a mesophilic,
gram-positive, non-motile, rod-shaped bacterium isolated on a mineral salts medium
containing 0.05% yeast extract, methanol, and acetate (Zeikus et al., 1980). The
morphology is pleomorphic, with single or branched rod structures predominating. The
importance of this organism was that it was the ﬁrst acidogenic bacterium which used
methanol as the sole carbon and energy source, producing a mixture of acetate and
butyrate during growth.

Several fundamental characteristics of the Marburg strain were determined by Zeikus
and coworkers in 1980, including a G+C content of approximately 49% and a very high
level of corrinoids when grown on methanol. The strain is also an obligate anaerobe, with
spore structures observed during heat shock treatment (Zeikus et al., 1980). Besides O2,
penicillin, streptomycin, tetracycline, and chloramphenicol were found to potent inhibitors
during growth on carbohydrates. Glucose, CH30H, CH30H and acetate, and H2 and
CO2 were all found to be usable energy sources, with a mixture of acetate and butyrate
being produced during growth except for H2 and CO2, where only acetate was produced
during growth (Zeikus et al., 1980). B. methylotrophicum is thus capable of growing

57

heterotrophically, autotrophically, or methylotrophically; abilities which provide it with an
extremely diverse range of possible substrates (Moench and Zeikus, 1983).

B. methylotrophicum was originally placed in the family Propionibacteriaceae, grouped
in the genus Butyribacterium after a similar organism, Butyribacterium reageri, which has
been isolated 40 years before (Barker, 1944). B. retrgeri, however, was later found to
actually be Eubacterium limosum, and thus some confusion, or contention, currently
exists (Sharak Genthner et al., 1981) as to whether B. methylotrophicum is also not a
separate genus but actually in the genus Eubacteria. The ongoing controversy has centered
on morphological and physiological characteristics such as spore-formation,
pleomorphism, vitamin requirements, and product spectrum. Interestingly, the high levels
of corrinoids found in B. methylotrophicum have raised the issue of whether this
organism is more suitable than other Propionibacteria for production of vitamin B12
(cobalamin), which is difﬁcult to synthesize chemically, although many species, including
B. methylotrophicum have been used for vitamin production (Hatanaka et al., 1988).

The CO strain of B. methylotrophicum was developed in 1981 by Lynd, Kerby and
Zeikus by ﬁrst growing cells of the Marburg strain on glucose, pyruvate, and CH30H in
the presence of a 100% CO gas headspace. Through repeated transfers on CH30H and
acetate in the presence of CO and then on 100% CO and 0.05% yeast extract without
alternate carbon sources, a culture was obtained which gradually adapted to use CO as the
sole carbon and energy some, which was subsequently designated the CO strain (Lynd et
al., 1982). The ability to use CO gas was retained through repeated back-transfers on
other substrates in the abscence of CO gas, thus conﬁrming the stability of the adapted
metabolism in the presence of selective pressure (Lynd et al., 1982).

13211 I'llll'Cl ..

Heterotrophic growth of B. methylotrophicum on substrates such as glucose proceeds
via the Emden-Meyerhof-Pamas pathway to produce pyruvate, which is then converted to

58

acetyl-CoA by pyruvate dehydrogenase. The overall products from glucose metabolism
are CO2, acetate, butyrate, H2, and cell mass, following the reaction,

C5H1205 ---> 0.31CO2 + 1.59CH3COOH + 0.21C3H7COOH + 0.13H2 + 1.58Cells
with a cell yield of 42.7 g cells (dry weight) per mole of glucose consumed (Lynd and
Zeikus, 1983). An empirical formula for B. methylotrophicum of C1H132Oo,5oNo_21
was also derived from an elemental analysis (Lynd and Zeikus, 1983). This carbon-
balanced reaction stoichiometry was calculated using carbon and available electron
balances (Erickson et al., 1978; Erickson and Oner, 1983), and was found to close within
0—1% for growth on glucose in a complex medium and to within 9-10% for growth on
glucose with a deﬁned medium (Lynd and Zeikus, 1983). These balancing methods are
detailed in the Appendix. A doubling time on glucose of 3-4 hr was calculated from the
batch growth data, and more than 25% of the consumed carbon was assimilated into cell
mass, as shown in the fermentation equation. Hydrogen production during glucose
metabolism is due to the activity of a hydrogenase enzyme, which gains electrons via an
NADH-linked ferredoxin (G. J. Shen and J. G. Zeikus, unpublished results). This
ferredoxin has been puriﬁed and its properties determined (Saeki et al., 1989a; Saeki et al,
1989b).

Activities for some of the major l-carbon compound metabolizing enzymes such as CO
dehydrogenase and formate dehydrogenase have been detected on glucose, pyruvate, or
lactate grown cells, indicating that these enzymes are still synthesized in the absence of CO
(G. J. Shen and J. G. Zeikus, unpublished results). The metabolic pathways for acetate
and butyrate production have also been elucidated during growth on glucose (G. J. Shen
and J. G. Zeikus, unpublished results), a pathway which also operates during growth on
CO and which is detailed in the next chapter. In general, the speciﬁc mechanims of carbon
and electron ﬂow have been investigated for heterotrophic metabolism of glucose, lactate,
and pyruvate, but at the time of this writing these results have not been completely

analyzed or published.

59

133! l'llll’Cl ..
Metabolism of l-carbon compounds in B. methylotrophicum has been elucidated both

in batch bottle experiments and in larger, laboratory scale batch and continuous
fermentations. The basic microbiology of this metabolism was determined by Zeikus and
coworkers in the early 1980's. The CO strain exhibits acetogenic metabolism, as described
previously, with formation of all cell mass and fermentation products arising from an
acetyl-CoA intermediate. The versatility of this organism is outlined Table 5 and is
described below.
Unicarbonotrophic metabolism of CO gas proceeds in batch culture by the carbon-
balanced reaction
4CO --> 2.17CO2 + 0.74CH3COOH + 0.45Cells
with an doubling time of approximately 12 hr. Carbon and electron recoveries during CO
metabolism were within 3%. Slight variations on this CO fermentation stoichiometry have
been observed, variations which include trace formation of butyrate during growth on C0
(Lynd and Zeikus, 1983). Activities for CO dehydrogenase and formate dehydrogenase
have been obtained during growth on CO, indicating that both of these enzymes participate
in unicarbonotrophic metabolism. Interestingly, the speciﬁc activity of CO dehydrogenase
is higher for CH30H and CO2 grown cells of the Marburg strain or the CO strain than for
CO grown cells of the CO strain, a strong reﬂection on the multipurpose nature of CO
dehydrogenase in acetogenic metabolism (Lynd et al., 1982; Kerby and Zeikus, 1987b).
Growth on H2 and CO2 produces acetate and cell mass with a 9 hr doubling time, and
exhibits the batch stoichiometry
2H2 + 1.03CO2 --> 0.43CH3COOH + 0.13Cells

with trace butyrate formation also observed (Lynd and Zeikus, 1983). Carbon and electron
recoveries were within 2-3% for this fermentation. As shown, the approximate H2:CO2

consumption ratio is 2:1 during growth. Consumption of H2 and CO2 in the presence of

60

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61

CO is fully inhibited, and only proceeds after a majority of the CO is consumed in cultures
incubated on C0, H2, and CO2 (Kerby and Zeikus, 1987a).
Growth of B. methylotrophicum with formate (HCOOH) as the sole substrate in batch

culture follows the stoichiometry

4HCOOH ---> 0.03H2 + 1.82CO2 + .95CH3COOH + 0.16Cells
with carbon and electron recovery within 3-5% and a 16 hr doubling time (Kerby and
Zeikus, 1987a). During this metabolism, H2 was ﬁrst produced and then later consumed,
furthermore, the overall growth on HCOOH compared to CO or H2 and CO2 was
signiﬁcantly reduced. However, cosubstrate fermentation of HCOOH plus H2 and CO2
yielded a carbon balance of

4HCOOH + 2.661-12 ---> 0.49CO2 + 1.66CH3COOH + 0.20Cells
with an elimination in the ﬂuxes for H2 production and consumption, and a carbon and
electron recovery within 10% (Kerby and Zeikus, 1987 a). Cosubstrate fermentation with
HCOOH and CO yielded a carbon balance of

4CO + 2.09HCOOH --> 2.9CO2 + 1.16CH3COOH + 0.52Cells
within a 6% recovery for carbon and electrons (Kerby and Zeikus, 1987a). The absence of
H2 production or consumption during growth on HCOOH and CO suggests that CO is a
strong inhibitor of the production hydrogenase enzyme, in that growth on HCOOH in the
absence of CO results in signiﬁcant H2 production (Kerby and Zeikus, 1987a). When
taken with the results for studies using CO, H2, and CO2, which suggest that CO is a
strong inhibitor of uptake hydrogenase activity, there is strong indirect evidence for
inhibition of both uptake and consumption hydrogenase by CO; however, this inhibition is
not complete, since H2 uptake was observed in the presence of low CO partial pressures
(Kerby and Zeikus, 1987a) and low hydrogenase activity was observed in CO grown cells
(Lynd et al., 1982). This inhibition may play a key role in the electron ﬂow during CO
metabolism, with ferredoxin oxidation proceeding via reduction of NADH as opposed to
via production hydrogenase and subsequent production of H2 gas (G. J. Shen, pers.

62

comm.). Determining the speciﬁc inhibition kinetics of CO on the hydrogenase enzyme is
thus an important and as yet uninvestigated area of unicarbonotrophic metabolism.

B. methylotrophicum is also capable of growth on CH30H in the presence of acetate,
with a batch carbon balance of

4CH3OH + 0.72CO2 + 0.04CH3COOH --->' 0.96C3H7COOH + 1.16Cells

with a 9 hr doubling time and a carbon and electron recovery within 5% (Lynd and
Zeikus, 1983). The principle of adding acetate during batch fermentation of CH30H in
this study was to regulate the acetate:butyrate product ratio; an increase in exogenous
acetate serves to decrease net acetate production and increase net butyrate production. This
stoichiometry is thus not the "true" carbon balance for methylotrophic metabolism. Such

experiments have also been conducted in the presence of acetate for cosubstrate

fermentations of CH30H and CO, and CH30H and HCOOH, with acetate produced in
the former and butyrate produced in the latter (Kerby et al., 1983). The ratio of CH30H to
CO2, or speciﬁcally to bicarbonate ion, is also a key parameter in regulating the
acetate:butyrate product ratio, with increasing concentrations of bicarbonate resulting in
increasing acetate production (Datta and O gletree, 1983). Presumably, these phenomena
are due to product inhibition, changes in the energetics of growth, and to changes in the
relative ﬂuxes of various pathways, all of which are subjects investigated in this study.

Results during extended batch CO fermentations with this organism have established
that fermentation pH directly inﬂuences the relative production of acetate and butyrate
during unicarbonotrophic growth on CO. At a pH of 6.8, acetate production was favored
over butyrate production with a molar ratio of approximately 32: 1, whereas at a pH of 6.0,
acetate and butyrate were formed in equimolar amounts (Worden et al., 1989).
Continuous, steady-state fermentations using B. methylotrophicum at several pH values
yielded similar results, with minor butanol and ethanol production occuring at lower pH
values (Grethlein et al., 1990). In this study, cell washout was observed at pH values
below 5.5.

63

The effect of fermentation pH on product formation was also investigated during
extended batch fermentation of C0, C02, and CH30H, and the same effect observed to a
lesser extent, with a 2-fold decrease in the acetate to butyrate ratio observed between pH
values of 6.8 and 6.0; however, acetate was still the primary product (Grethlein, 1989). A
preliminary investigation as to the overall effect of fermentation pH on key catabolic
enzymes was also initiated, and a 22% increase in the speciﬁc activity of butyrate kinase
observed between pH values of 6.8 and 6.0 (Grethlein, 1989). Although not by any
means a conclusive result, this preliminary enzymatic study suggested that the overall
effect of fermentation pH in regulation of CO metabolism includes regulation of speciﬁc
enzyme activities. Conﬁrmation of this hypothesis has since been obtained for the key
enzymes involved in acetate, butyrate, ethanol, and butanol formation during CO

metabolism (J. S. Sheih, unpublished results).

1.4. Biological Processing of Coal-Derived Gases

The use of coal-derived synthesis gas as an industrial feedstock for production of fuels
and chemicals has provided an increasingly attractive alternative to present petroleum-
based processes stemming from recent developments in C1 chemistry and coal gasiﬁcation
technology. One of the major drawbacks in using coal-derived synthesis gas is the
required stringent removal of catalyst poisons such as H2S , C08, and other trace
contaminants. The levels of these sulfur contaminants can vary between 0-2% of the
synthesis gas. Removal of these sulfur species requires energy-intensive puriﬁcation steps
which add signiﬁcantly to the cost of the ﬁnal product, particularly when the high sulfm'
content of most US. coals is recognized (Wilson et al., 1988).

Biological processing of synthesis gas, a potential option to conventional catalytic
processing, is focused on a reduction in the amount of puriﬁcation and synthesis gas
preparation steps required for efﬁcient conversion to fuels and chemicals. Research in this

area is presently in the exploratory phase and is focused on characterization and

64

development of microorganisms and anaerobic fermentations for conversion of synthesis
gas components. Some of the key motivations and obstacles in a biological processing
approach, as well as a process concept for synthesis gas bioconversion, are described in
this section.

HIE .1“ {3.1.12 .

The main advantage to be gained by employing a biological route for coal-derived
synthesis gas conversion is in reduced requirements for sulfur gas removal and clean-up
prior to conversion. The sulfur, or more speciﬁcally sulﬁde, tolerance of anaerobic
microorganisms capable of CO and H2/CO2 metabolism has been documented as at or
above a concentration 2% H28 or COS (Vega et al., 1990; Grethlein et al., 1991a). With
sulﬁde tolerant microorganisms, raw or unpuriﬁed synthesis gas can potentially serve as
the feedstock for fermentation, without the intensive sulfur puriﬁcation that is typically a
major part of conventional synthesis gas processing. Operating costs for production of
clean synthesis gas, including the coal preparation, gasiﬁcation, and gas puriﬁcation, can
account for over half of the total operating costs for coal-based processes such as at
SASOL (Dry, 1987). The potential savings in a biological route will thus be realized in
reduced capital and operating costs for puriﬁcation equipment and energy requirements.

The other major advantage of biological processing is in the area of synthesis gas
composition. The speciﬁcity of chemical reactions based on reactant stoichiometry requires
set CO/H2 ratios in the synthesis gas prior to catalytic reaction. Obtaining this set
composition requires separate gas-shifting reactors, which frequently operate under high
pressure, in order to modify the CO/H2 ratio via the water-gas shift reaction. These shift-
conversion operations often require gas recompression, high temperatures, and reactor
equipment. In biological fermentations, a set CO/H2 ratio is not of importance with respect
to the metabolic reactions. Substantial savings in both operating and capital equipment
costs could thus be realized by the removal of gas-shifting operations from the synthesis

gas conversion process.

65

Fermentations at relatively low temperature and pressure will bring advantages in
operating costs for the actual synthesis gas reactors, temperatures and pressures which are
generally at least an order of magnitude less than those required in catalytic reactors using
synthesis gas as a feedstock. The efﬁciency of anaerobic fermentations, the speciﬁcity of
most biological reactions, and the high product yields of anaerobic bacteria are also
speciﬁc advantages of biological routes in comparison to typical chemical processing.

“22 °lD° l [E'] . IE .

When comparing biological routes for synthesis gas conversion to chemical or catalytic
routes, the most prevalent disadvantages are in the achievable reaction rates, the product
separations costs based on typically aqueous, dilute fermentation broths, and the issues of
culttne maintenance, stability, and sterility. The intrinsic biological reaction rates for CO or
H2 vary greatly depending on the microorganism, but the achievable reaction rates will
depend on the mass-transfer characteristics of aqueous fermentations, based on sparingly
soluble gases such as CO and H2. Achieving high mass transfer rates could potentially
require high energy input and thus incm high operating costs.

Separation of aqueous fermentation products, particularly alcohols, typically requires
energy intensive distillation unit operations, for which the operating costs increase as the
concentration of alcohol decreases. There is also a practical minimum limit for separating
ethanol or butanol from water, generally in the range of 10-15 g/L. The productivity and
product concentrations resulting from a bioconversion process based on coal-derived
synthesis gas will thus be critical in the assessment of its potential. Fermentation and
process sterility is also a signiﬁcant concern when working with biological systems;
contaminated systems can result in uncontrolled reactions, decreases in product yields, and
a potential for the desired organism to be out-competed for substrates and nutrients.

Development of fundamental scientiﬁc knowledge and fermentation technology for the
types of unique biological systems which would potentially be employed for processing of
coal-derived synthesis gas is thus a challenging endeavor from both a biological and an

66

engineering standpoint. None of the potential disadvantages associated with biological
systems in this context are so overwhelming as to discount the entire concept, since
separations, reaction control, and process stability issues will be present regardless of
whether a biological or chemical route is used. However, the issues of reaction rate and
product concentration are probably the most pervasive when discussing the limitations of a
biological route. The advantages of biological systems are much more distinguishing in
that all catalytic systems presently known are highly susceptible to sulﬁde poisoning,
whereas in biological systems, particlularly for anaerobic bacteria, many species are
known to be sulﬁde tolerant. The degree of tolerance, however, varies from organism to
organism. Similarly, regulation of the water-gas shift reaction as a prelude to chemical
processing is an integral aspect of all current chemical routes, whereas, in general, the
synthesis gas composition requirements for biological routes can be met in the gasiﬁer
itself, and will not be a function of the microorganism. These features have thus prompted
government funding of two exploratory research efforts, at the University of Arkansas and
at the Michigan Biotechnology Institute, for investigation of the potential for biological
conversion of coal-derived synthesis gas.

II3II “EH 1. 1' E . E

With support from the US Department of Energy, and more recently by the Center for
Research on Sulfur in Coal and the Illinois Coal Development Board, the Michigan
Biotechnology Institute (MBI), in Lansing, MI, has been developing a bioconversion
process for coal-derived synthesis gas, of which the present work is an integral part. This
process is aimed at biological production of fuel oxygenates, speciﬁcally alcohols, from
the CO and H2 in synthesis gas. The overall goals of the MBI process are to establish a
fundamental knowledge base for high productivity, continuous, anaerobic fermentations
based on C0 and based on the possible fermentation products from C0.

The original MBI process scheme consisted of a two-stage anaerobic fermentation of

coal-derived synthesis gas plus methanol to produce an aqueous mixture of butanol,

67

ethanol, and acetone, followed by energy efﬁcient distillation recovery. This process
scheme is outline in Figure 8, where the synthesis gas mixture of CO and H2, plus
methanol, enters the ﬁrst stage, which is an acidogenic fermentation based on C0 gas.
This fermentation produces a mixture of acetic and butyric acids, which are then
transferred to the second stage where a solventogenic fermentation converts the acids to
the corresponding alcohols ethanol and butanol, plus minor amounts of acetone. The
chosen organisms for these fermentation stages were the CO strain of B .
methylotrophicum for the acidogenic stage, and a mutant strain of Clostridium
acetobutylicum for the solventogenic stage. The hydrogen gas is presumably consumed in
the second stage fermentation to provide electrons for solventogenesis. Cell recyling is
used for maintaining high cell density fermentations in both stages of the process.
The overall theoretical conversion reactions in the acidogenic stage are
4CO + 2H20 ---> CH3COOH + 2CO2
CH30H + 0.2CO2 ---> 0.3C3H7COOH + 0.8H20
where methanol is introduced to promote butyrate formation. The exit streams from the
ﬁrst stage will thus be gaseous H2 and an aqueous mixttn'e of acetic and butyric acids. The
acids then enter the second, solventogenic stage, where the theoretical reactions
0.7CH3COOH + 1.4H2 ---> 0.7C2H50H + 0.7H2O
0.3CH3COOH ---> 0.15CH3COCH3 + 0.15CO2 + 0.15H2O
0.3C3H7COOH + 0.6H2O ---> 0.3C4H90H + 0.3H2O
result in mixture of butanol, ethanol, and acetone. The overall process theoretical

fermentation balance is thus
4CO + CH30H + 2H2 + 0.05H2O --->

0.3C4HgOH + 0.7C2H50H + 0.15CI-I3COCH3 + 1.95CO2
With an improved mutant adapted for decreased acetone formation, that product can

potentially be eliminated from the solventogenic fermentation step.

68

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The proposed energy efﬁcient solvent recovery scheme comprises a patented distillation
process (II-IOSR), based on heat recovery via vapor recompression and optimal side-
stream return (Grethlein and Lynd, 1986). This distillation technology has exhibited
dramatic energy savings for binary ethanol—water systems, and is expected to offer similar
savings in a multicomponent mixture such as for the MBI bioconversion process. The use
of pervaporation technology directly applied to the solventogenic fermentation has also
been proposed. Overall, some of the most unique aspects of the MBI bioconversion
concept include direct conversion of raw synthesis gas, conversion of high CO/I-12 ratio
synthesis gases to liquid oxygenates, and an integrated, energy-efﬁcient product
separation.

Since its inception in 1987, the MBI process has undergone a signiﬁcant amount of
development and changes based on fermentation and microbiology research for both the
acidogenic and solventogenic stages. As described previously, The CO strain of B.
methylotrophicum has exhibited butyrate formation from C0 (Worden et al., 1989), as
well as a capacity for butanol and ethanol production directly from CO gas. However,
alcohol production was observed only in trace quantities less than 0.1 g/L (Grethlein et al.,
1990). A continuous fermentation for CO gas was developed, and operation at long-term,
steady-state conditions was achieved (Grethlein et al., 1990). Fermentation pH was found
to be a key regulatory factor in the overall CO metabolism of this organism during these
studies. Other key developments include application of cell recycling to the solventogenic
fermentation, with a 3 fold increase in the rate of solvent production observed. Uptake of
butyrate by C. acctoburylicwn has also been demonstrawd under a variety of conditions.

There were several key issues for the biocoversion process upon which the initial
objectives of the exploratory research effort were based; issues which still had signiﬁcant
bearing at the beginning of this work. These were,

1. Low rates of product formation and growth by anaerobic, CO-utilizing bacteria.

2. Low rates of organic acid uptake by solventogenic bacteria.

70
3. Difﬁculties of long-term, steady-state maintenance of anaerobic fermentations.

4. Low solubilities and mass transfer limitations of CO and H2.

5. Product inhibition by alcohols and acids during anerobic fermentation.

6. High energy requirements for separation of dilute aqueous solvent mixtures.

These issues have been important in focusing MBI's research efforts towards obtaining
fundamental knowledge of the particular microbial systems, as well advancing the
technical aspects of the bioconversion process concept.

1110].. 15 “E IN]

The work in this dissertation is focused on the research and development of the ﬁrst
stage, acidogenic CO fermentation, which uses the CO strain of B. methylotrophicum.
Based on previous results from batch and chemostat systems, the overall objectives for the
fermentation and process development were:

1. Increase CO fermentation productivity and efﬂuent product concentrations.

2. Operate continuously at low fermentation pH values, where butyrate and

alcohol production are favored.

3. Regulate CO metabolism away from acetate and towards other products.

4. Develop a liquid media capable of supporting high cell densities.

In conjunction with these development objectives were several research objectives:

1. Elucidate pathways for alcohol production from CO in B. methylotrophicum.

2. Investigate effect of fermentation products on C0 metabolism.

3. Identify possible mechanism(s) of metabolic regulation during CO metabolism.

4. Determine the organism's tolerance for sulﬁde species that may be present in coal-

derived synthesis gas .

5. Determine the overall mechanism of the pH effect on CO metabolism.

The research effort is thus concerned both with the technical development of the CO
fermentation as well as with increasing the scientiﬁc knowledge of CO metabolism and its

regulation. The challenge of this system is particularly apparent in its novelty, for there

71

were no previous examples of similar studies with CO fermentations to be used as
guidelines for the present work. The ultimate goal was to determine the potential for a
biological fermentation based on CO from coal-derived synthesis gas, and to obtain a
fundamental understanding of this speciﬁc biological system. Such an understanding
allows identiﬁcation of the important parameters and limitations, regardless of the speciﬁc

microorganism or fermenter conﬁguration used.

CHAPTER 2

EXPERIMENTAL MATERIALS AND METHODS

2.1. Microorganism and Culture Conditions

The CO strain of B. methylotrophicum was originally obtained from a frozen stock
culture at the Michigan Biotechnology Institute (courtesy of Dr. M. K. Jain) in Lansing,
Michigan, USA. A stock culture was maintained in 152 mL serum bottles (Wheaton,
Millville, NJ) sealed with butyl rubber bungs (Bellco Biotechnology, Vineland, NJ) and
aluminum crimp caps (Bellco Biotechnology, Vineland, NJ). These bottles contained 50
mL of phosphate buffered medium under an initial gas headspace of 100% CO gas at
approximately 2.1 atrn total pressure. When applicable, 28 mL pressure tubes (Bellco
Biotechnology, Vineland, NJ) were substituted for the bottles, containing 10 mL of liquid
media. The stock culture was grown in the dark at 37°C on a shaking platform (New
Brunswick G10 Gyrotory Shaker, New Brunswick Scientiﬁc Co., Edison, NJ) at 100
rpm. The culture were transferred to fresh bottles every 2-3 weeks. Periodically, cells
from a fermentation experiment were substituted for transferred cells in the continuation of
the stock culture, so that any cellular adaptations resulting from long-term fermentation
could be maintained. Twice during the course of this work, a complete set of 5 mL
glycerol containing cell cultrne bottles was frozen at -70°C for extended storage of the cell
line. Fermentation studies were initiated by inoculation with a 0.5% - 2.0% (v/v) actively

growing B. methylotrophicum stock culture.

2.2. Culture Media and Gases
A phosphate buffered basal medium was adapted from Lynd et al., (1982), and used

for maintenance of the stock culture as well as in all bottle experiments and batch

72

73

fermentation experiments. The pre-autoclaving components of this media are shown in

Table 6. Trace Mineral II solution is a aqueous mineral salts addition containing 12.8 g/L

Table 6. Components of the phosphate buffered basal medium.

 

 

Component Amount (per Liter)
Double Distilled water 1 L

NaCl 0.9 g
MgCl202H2O 0.2 g
CaClz-2H2O 0.1 g
NH4Cl 1.0 g
Trace Mineral II 10 mL
0.2% Resazurin 1.0 mL

 

nitrilotriacetic acid, 0.10 g/L FeSO4-7H20, 0.10 g/L MnCl2-4H20, 0.17 g/L
CoC12°6I-I20, 0.10 g/L CaCl2-2H20, 0.10 g/L ZnCl2, 0.02 g/L CuCl2-2H20, 0.01 g/L
H3BO3, 0.01 g/L sodium molybdate, 1.0 g/L NaCl, 0.017 g/L Na2SeO3, and 0.026 g/L
NiSO4°6H20. The nitrilotroacetic acid is not a carbon source, but serves to prevent
precipitaion of iron in the form of FeS. Resazurin is a colorometric indicator of dissolved
oxygen in the pH range of 58

After autoclaving, 10 leL of a 10% w/v DIFCO bacto yeast extract (Difco
laboratories, Detroit, MI) solution and 25mIJL of an aqueous phosphate buffer containing
150 g/L KH2PO4 and 290 g/L K2HPO4 were added. Also added was 10 M of an
aqueous vitamin solution containing 0.002 g/L Biotin, 0.002 g/L Folic acid, 0.010 g/L
B6¢HC1 (pyridoxine), 0.005 g/L B 1-HC1 (thiamine), 0.005 g/L B2 (riboﬂavin), 0.005 g/L

74

nicotinic acid, 0.005 g/L pantotlrenic acid, 0.0001 g/L crystalline B12 (cyanocobalamin),
0.005 g/L PABA (para-aminobenzoic acid). The ﬁnal addition was 25 M of a 2.5%
w/v Na2S°9H2O aqueous reducing aqueous solution. All the post-autoclaving additions
above were sterilized seperately, either by autoclavin g or ﬁlter sterilization. This media
thus consists of mostly inorganic salts, minerals, yeast extract, and vitamins, and contains
no carbon source except for the small amount present in the yeast extract. The continuous
fermentation experiments were conducted with the above media but with 10 mIJL

phosphate buffer added instead of 25 leL, and with 10 M of an aqueous cysteine-
sulﬁde reducing solution containing 12.5 g/L cysteine sulﬁde and 12.5 g/L Na28-9H2O

substituted for the 25 M Na2S solution. The continuous cell recycling fermentations
were conducted with a 2-fold increase in the amount of yeast extract added, a 5-fold
increase in the amount of Trace Mineral II solution added, and a 10-fold increase in the
amount of vitamin solution. These fermentations also used 1.98 g/L (NT-I4)2SO4 and
10mL/L of a modiﬁed ammonium phosphate buffer containing 150 g/L KI-12PO4, 116 g/L
K2I-IPO4, and 132 g/L (NH4)2HPO4 in double distilled water.

The CO, CO2, N2, and C08 gases used in this study were obtained either from Linde
(Linde Division, Union Carbide, Co., Warren, MI), Airco (Airco Industrial Gases, Royal
Oak, MI), or from AGA (AGA Gas Inc., Cleveland, OH). Gas puriﬁes averaged 99.0%
for CO, 99.99% for CO2, 99.99% for N2, and 98.5% for COS. H2S gas was generated
by adding a Na2S solution to an excess of HCl under a vacuum headspace. The

approximate purity of this synthesized H2S gas was 90%.

2.3. Fermentation Equipment

All fermentation systems were designed for pure culture operation in an anaerobic
environment. Nitrogen gas was used to initially ﬂush the systems prior to inoculation.
Substrate gases were passed through a hydrogen-reduced copper catalyst oven (Sargent
Welch Scientiﬁc Co., Skokie IL) to remove any trace oxygen from the gas. Sample ports

75

were stoppered with butyl rubber bungs and crimped aluminum caps. All of the
fermentation systems typically used Masterﬂex food grade Norprene tubing (Cole Parmer
Instrument Co., Chicago, IL) for gas and liquid lines, and sealed connections with
Dennison 7.5 inch Bar-10k cable ties (Dennison Manufacturing Co., Fastener Division,
Framingham, MA). Fermentation vessels and attached external parts were autoclaved for a
minimum of 30 minutes, and up to 95 minutes for the continuous systems with large
liquid media reservoir volumes. Batch systems were autoclaved with the liquid media
already in the fermenter, continuous systems were autoclaved separately from the media.
External CO gas concentration was monitored for safety using both an SMC Series 2006
Carbon Monoxide Monitor and an SMC Series 2001 Combustible Gas Monitor (Sierra
Monitor Corporation, Millipitas, CA), with both sensors calibrated in the toxic range of 50
ppm CO gas.

The continuous cell recycle fermentations were conducted in 2 L Multigen fermenters
(New Brunswick Scientiﬁc Co., Edison, NJ, USA) with working volumes of either 1.2
or 1.4 L. A New Brunswick TM-100 Series Mega-Flow Filtration System (New
Brunswick Scientiﬁc Co., Edison, NJ, USA) using an Amicon YM30 0.3 pm pore size
cellulosic membrane (Amicon Corporation, Danvers, MA, USA) was used for cell
microﬁltration. Gas recycling was conducted using a Cole Parmer Masterﬂex peristaltic
pump (Cole Parmer Instrument Co., Chicago, IL). A porous, stainless steel frit was used
to reduce gas bubble size into the liquid from this gas recycle stream. Fermentation pH
was automatically controlled by addition of 2N NaOH solution, using an Ingold Type 465
pH electrode (Ingold Electrodes Inc., Wilmington, MA) in conjunction with either a New
Brunswick Model pH-40 Automatic pH controller (New Brunswick Scientiﬁc Co.,
Edison, NJ) or a Chemcadet pH Meter/Controller (Cole Parmer Instrument Co., Chicago,
IL). Liquid ﬂows were established using Gilson Minipuls 2 peristaltic pumps (Gilson
Electronics, Middleton, WI). The liquid media reservoirs were prepared in 16 L batches
using 20 L PYREX solution bottles (Corning, Inc., Corning, NY). Outlet gas ﬂow rates

76

were measured using Cole Parmer series N062-01 or Series N032—41 rotameters (Cole

Parmer Instrument Co., Chicago, IL).

2.4. Fermentation Conditions

In general, bottle experiments were conducted on a shaking platform in the dark as
described previously for the stock culture maintenance. Bottles were initially at a pH of
between 7.0 and 7.1 with a 100 % CO gas atmosphere. The pH was not controlled in
these systems. Gas phase components such as H2S gas were added by volume to the
bottles prior to inoculation. Liquid phase components such as organic acids and alcohols
were also added by volume from stock solutions prior to inoculation. Both gas and liquid
samples were taken on a daily basis in these systems. Test tube growth experiments
measured optical density by direct insertion into the spectrophotometer, and did not
involve any external assays.

The experiments involving acids and alcohols were conducted in 152 mL sealed,
anaerobic bottles with 50 mL liquid media and an initial 100% CO gas headspace at 2-2.2
atm, thus achieving a 2:1 gas to liquid volume ratio. Initial concentration ranges of
approximately 0- 12 g/L for alcohols were achieved by volumetric addition of either ethanol
or n-butanol directly to the anaerobic bottles. Initial concentration ranges of approximately
010 g/L for acetate or butyrate were achieved by volumetric addition of stock solutions
containing either 4 M acetic acid or 3 M butyric acid. The broth pH was adjusted prior to
inoculation during organic acid tolerance experiments by equimolar volumetric addition of
10 N NaOH solution. To determine the effect of sodium on the growth of B .
methylotrophicum, 28 mL sealed, anaerobic culture tubes, each with 10 mL of phosphate-
buffered media under an 18 mL 100% CO gas headspace at 1.5-1.6 atm total pressure
were used. A concentration range of 0 to 0.5 M sodium was generated by volumetric
addition of sodium from a stock 5 M NaCl solution. All growth phase tolerance
experiments were initiated by the addition of a 2% v/v mid-log phase B. methylotrophicum

77

stock culture. All bottle experiments were conducted in duplicate. Samples for cell density,
broth pH, CO gas, CO2 gas, and all liquid phase products were taken on a daily basis.

For experiments testing sulﬁde tolerance, varying sulﬁde concentrations were generated
by directly adding H2S gas or by addition of a stock 2.5 % Na2S solution to 152 mL
anaerobic bottles containing 50 mL of phosphate-buffered media with a 100% CO gas
headspace at 2-2.2 atm. Initial equilibrium concentrations of the total liquid sulﬁde species
were calculated for each experiment. Control experiments were always conducted with the
standard total sulﬁde concentration of 0.083 g/L (based on the pre-equilibration
concentration). All bottle experiments were conducted in duplicate. Experiments were
initiated by the addition of a 2% v/v mid-log phase B. methylotrophicum stock culture.
Samples for cell density, pH, CO gas, C02 gas, and all liquid phase products were taken
on a daily basis.

The continuous cell recycle fermentations had a CO gas addition rate of 40 mL/min,
with a gas recycling rate from the headspace into the liquid broth at 150 mL/min.
Fermentation parameters for these experiments included a dilution rate of 0.015 hr'l, an
impeller speed of 450 rpm, a temperature of 37°C. This speciﬁc system is diagrammed in
detail in Chapter 4.

2.5. Fermentation Broth Analysis

Liquid samples were taken using N2-ﬂushed, sterile syringes inserted into the butyl
rubber bungs in either the bottles or fermenter sample ports. Approximately 1-5 mL liquid
volume was taken, depending on the experiment. Gas volume samples were taken using
gas tight glass syringes with a sample volume of 1.0 mL. Cell density was measured by
determining optical density at 660 nm on a Sequoia-Turner Model 340 spectrophotometer
(Sequoia-Turner Corporation, Mountain View, CA). Samples with an optical density
greater than 0.6 were diluted 10 fold with distilled water. Those with an optical density
greater than 6.0 were diluted 100 fold. Cell mass determinations for B. methylotrophicum

78

were made using a previously derived dry weight versus optical density calibration curve
(Lynd et al., 1982). This curve was periodically reproduced in order to ensure accuracy
for the cell mass calculations based on a value of 0.344 g cell mass/L-OD (Lynd et al.,
1982). Undiluted samples were observed microscopically for general viability and
contamination and subsequently tranferred to 1.5 mL Eppendorf polyproplyene micro test
tubes (Brinkmann Instruments, Inc., Westbury, NY), centrifuged for 2 minutes at 12,000
rpm in an Eppendorf Centrifuge 5415 (Brinkmann Instruments, Inc., Westbury, NY),
and the pH checked to provide external calibration for the experiments.

2.6. Fermentation Product Analysis

Concentrations of organic acids and alcohols were determined on a Hewlett-Packard
5890A Gas Chromatograph in tandem with a 3392A auto-sampler and 3392A integrator
(Hewlett—Packard Co., Avon, PA), equipped with a 1.2 m Chromosorb 101 80/100 mesh
column and a ﬂame-ionization detector. The N2 carrier gas ﬂow rate was 30 mL/min.
Operating temperatures were 190°, 220°, and 250°C for the column, injector, and
detector, respectively. All samples were centrifuged, acidiﬁed with 3N phosphoric acid (1
part acid / 10 parts sample) and then transferred to sealed injection vials. All samples were
automatically injected in duplicate using the autosampler. Samples were compared to 10
mM standards run concurrently.

Analysis of organic acids and alcohols was alternately conducted on a Hewlett-Packard
5890A Gas Chromatograph in tandem with a 7673A auto-sampler and 3392A integrator
(Hewlett-Packard Co., Avon, PA), equipped with a 1.2 m Carbopack C 80/100 mesh
column and a ﬂame-ionization detector. The N2 carrier gas ﬂow rate was 15 mL/min.
Operating temperatures were 125°, 150°, and 220°C for the column, injector, and
detector, respectively. All samples were centrifuged, acidiﬁed with 20 % (v/v) formic acid
(1 part acid / 10 parts sample) and then transferred to sealed injection vials. All samples

79

were automatically injected in duplicate using the autosampler. Samples were compared to
10 mM standards run concurrently.

The analysis of butanol and ethanol was conﬁrmed periodically by high-performance
liquid chromatography using a Waters 410 differential reﬁactometer (Waters Instruments,
Division of Millipore, Milford, MA, USA) with an Aminex ion exclusion HPX 87H
column at 40°C, and in an outside laboratory by gas chromatography-mass spectroscopy
using a Hewlett-Packard 5840 gas chromatograph and Hewlett-Packard 5895 Quadrapole
Mass Spectrometer (Hewlett-Packard Co., Avondale, PA, USA) with a Chromosorb 101
80/ 100 mesh column at 190°C.

CO, CO2, N2 and H28 gases were analyzed by gas chromatography in a GowMac
Series 580 Gas Chromatograph (GowMac Instrument Co., Bridgewater, NJ) using a 1.8
m Carbosphere 80/ 100 mesh column and a thermal conductivity detector. Column,
injector, and detector temperatures were 100°, 115°, and 150°C, respectively. A helium
carrier gas was used with a ﬂow rate of 50 mL/min. The detector current was 200 mA. All
injections were 0.4 mL volume and at atmospheric pressure. Samples were compared to a

calibrated standard curve constructed from pure gas responses.

CHAPTER 3

METABOLIC TOLERANCE AND REGULATION STUDIES

3.1. Objectives

The primary objective of the research described in this chapter was to investigate the
effects of the known CO fermentation products on CO metabolism in B .
methylotrophicum. These questions arose from the results obtained during the continuous
cell recycling fermentation experiments described in Chapter 4, and thus chronologically
these experiments would actually follow those results. The key objectives of these
metabolic regulation and tolerance experiments were to determine the potential for alcohol
or acid inhibition, the potential for acid uptake, the overall effect, if any, of these
compounds on C0 metabolism, and also, from a processing standpoint, the limits of
sulﬁde tolerance. Also, an enhanced media was developed for use in the high cell density
fermentations described in Chapter 4. These microbiological studies are presented prior to
the fermentation studies, to allow a complete discussion of the results in Chapter 4.

Overall, the majority of the experiments described in this chapter were focused on
elucidating mechanisms for manipulation or regulation of product formation during CO
metabolism. The choice of an experimental system for these studies was thus critical, since
the ultimate application of the information obtained was towards analyzing fermenter
results. The impracticality of conducting controlled variable, microbiological experiments
in a lab-scale set-up, particularly for an organism with a 12 hr doubling time, prompted the
choice of batch bottles as an experimental sytem for these experiments. The advantages of
this set-up include the ability to conduct experiments with controls and experimentals in
duplicate or triplicate runs, high practicality in the number of experiments possible at one
time, and in the short-term duration of the experiments. However, several limitations of

the batch bottles must be recognized, particularly the difﬁculties in controlling only one
80

81

variable, and in extrapolating batch growth results for analysis of continuous fermenter
data. Furthermore, the limits of the bottle system allowed only growth phase experiments
to be conducted, and this origin must be considered when analyzing stationary phase

fermenter results.

3.2. Effect of Alcohols on Growth and Product Formation

Although B. methylotrophicum produced butanol and ethanol from C0 in
concentrations ranging from 0.005 to 0.01 g/L during chemostat operation (Grethlein et
al., 1990), concentrations which are not likely to be inhibitory either to growth or to
product formation, the possibility during a high productivity CO fermentation for alcohol
or even acid inhibition deﬁnitely exists. There is a strong potential for gradual adaptation
of this organism towards alcohol production during prolonged, low pH fermentation of
CO, and such adaptation would be limited, in part, by any potential product inhibition
effects. A 1020 fold increase (an optimistic but not unrealistic goal) in the alcohol
concentration during a cell recycle fermentation could inﬂuence the cellular metabolism.
Alcohol tolerance is, therefore, a potentially rate-limiting aspect of high cell density CO
fermentations. Thus, in order to investigate any possible inhibition by butanol and ethanol
during CO fermentation, several experiments were conducted and the rates of cell growth
and product formation were measured both with and without the presence of butanol or
ethanol. The results from these studies have a direct impact on the feasability of any acid
or alcohol-producing fermentation for B. methylotrophicum using CO gas as the sole
carbon and energy source.

The effect of ethanol and butanol on the fermentation pattern during growth of B.
methylotrophicum using CO as the sole carbon and energy source was determined in
batch, anaerobic bottles. Time-course growth curves for the culture optical density as a
function of alcohol concentration between 0-12 g/L are shown in Figures 9 and 10 for the

ethanol and butanol experiments, respectively. Carbon and available electron balances

Opdcal Density (660 nm)

Optical Density (660 mu)

 

 

 

 

 

 

 

 

1.60 I a I a I a I r I a I a
1.40 C-
1.20 E-
1.00 _-
0.80 _-
0.60 -
I
0.40 _-
0.20 -

0.00 I I T I I r i ] 1 f I
0 24 48 72 96 120 144

Time (hrs)

Figure 9. Effect of increasing ethanol concentration during
growth of B. methylotrophicum on C0 gas.

 

 

 

 

 

 

1.60 I a I 1 I r I r I 1 I r I 1 I r I r I 1
+Control 0* -_-
1.40 +1.1 ”L l:
1.20 . u“ E-
..._s.4ur. .
1.00 +10-3uL :-
0.80 E.
0.60 '-
t
0.40 :
l-
0.20 r
0.00 “:2— ___ — ‘i— .-
0 24 48 72 96 120 144 168 192 216 240

Time (hrs)

Figure 10. Eﬂ'ect of increasing butanol concentration during
growth of B. methylotrophicum on C0 gas.

83

(Erickson, et al., 197 8; Erickson and Oner, 1983), calculated assuming 100%
conservation during these experiments, are shown in Tables 7 and 8. These balances have
been shown to be accurate to within 3-5% in similar experimental systems using this
organism grown on C0 gas (Lynd et al., 1983). Table 9 shows several observed and
calculated variables as a function of alcohol concentration, including the maximum cell
density obtained as a function of the control experiment (no exogenous alcohol added),
cell density, the total drop in pH, the ﬁnal acetate and butyrate concentrations observed,
and the corresponding acetate to butyrate concentration ratio. Table 10 shows the observed
doubling time during mid-log phase and also the total consumption of CO gas. For
ethanol, concentrations of 5 g/L and above signiﬁcantly inhibit both growth and product
formation. For butanol, the tolerance level is somewhat lower, with signiﬁcant inhibition
of cell growth occuring at butanol concentrations of 2.8 g/L and higher.

The same effect over the concentration range of 0-2.5 g/L was examined in pressure
tubes, with time-course measurement of optical density during experiments with either
ethanol or butanol. Results for these experiments are shown in Figures 11 and Figure 12.
These data show no signiﬁcant effect of either alcohol between 0-l.0 g/L on the rates of
cell growth, but with some decrease in the growth rate at concentrations of 2.5 g/L for
both alcohols, an effect more signiﬁcant for butanol, where a marked decrease in the ﬁnal
cell density is also observed, relative to the other experimentals and to the contol. These
ﬁndings are consistent with the trends observed in the previous bottle experiments,
showing butanol to be somewhat more toxic than ethanol to cell growth.

The stationary phase response to exogenous ethanol and butanol was also examined.
Fermenter grown cells were transferred at the onset of the stationary phase to batch bottles
and incubated under CO gas headspaces in the presence of 7.1 g/L ethanol or 5.3 g/L
butanol. These expriments were inoculated from separate fermentation cultures. Table 11
shows the carbon and available electron balanced fermentation stoichiometries calculated

over these 3 day experiments. Initial concentrations of acetate and cell mass were 3.9 and

84

Table 7. Balanced carbon and available electron equations for batch incubation in the
presence of ethanol during growth on C0 gas.

 

 

Ethanol Concentration.
Fermentation Stoichiometr'ya
0.0 g/L 4CO --> 2.16CO2 + .48AC + .10BU + .03EH + .40CM
1.4 g/L 4CO --> 2.17CO2 + .45AC + .13BU + .02EH + .36CM
2.8 g/L 4CO --> 2.18CO2 + .43AC + .15BU + .01EH + .35CM
5.4 g/L 4CO --> 2.07CO2 + .71AC + .05BU + .31CM
11.0 g/L 4C0 -> 2.01CO2 + .87AC + .25CM

 

aSymbols used: AC = Acetate; BU = Butyrate; EH = Ethanol; CM = Dry Cell Mass.

Table 8. Balanced carbon and available electron equations for batch incubation in the
presence of butanol during growth on CO gas.

 

 

Butanol Concentration.
(8’14) Fermentation Stoichiometr'ya
0.0 g/L 4CO --> 2.16CO2 + .42AC + .12BU + .01EH + .47CM
1.1 g/L 4CO --> 2.19CO2 + .36AC + .14BU + .02EH + .48CM
2.8 g/L 4CO --> 2.15CO2 + .51AC + .04BU + .09EH + .SOCM
5.4 g/L 4CO -> 2.03C02 + .73AC + .52CM
10.3 g/L 4CO -—> 2.00CO2 + .99AC + .03CM

 

aSymbols used: AC = Acetate; BU = Butyrate; EH = Ethanol; CM = Dry Cell Mass.

85

Table 9. Effect of increasing alcohol concentration on cell density and product formation
during growth on C0 gas.

 

Alcohol% Maximum OD pH Change Final Concentration (mM) Product Ratio

 

Concentration Obtained (ApH) Acetate Butyrate (AczBu)
Ethanol:

Control 100 0.94 23.8 4.8 5.0
1.4 g/L 86.7 0.92 21.8 6.5 3.4
2.8 g/L 78.0 0.84 19.4 6.6 2.9
5.4 g/L 37.9 0.65 17.1 1.3 13.2

11.0 g/L 17.0 0.28 10.5 0.0 cc

Butanol.
Control 100 0.97 19.0 5.6 3.4
1.1 g/L 98.9 0.90 15.8 6.1 2.6
2.8 g/L 43.0 0.47 9.0 0.7 12.9
5.4 g/L 9.0 0.09 2.3 0.0 cc
10.3 g/L 2.0 0.00 0.5 0.0 co

 

Table 10. Effect of increasing alcohol concentration on cell growth and CO consumption

 

 

during growth on C0 gas.

Alcohol % Maximum OD Total 00 Consumption Observed Doubling
Concentration Obtained (umoles) Timea (hrs)
Ethanol.

Control 100 10,500 12.8
1.4 g/L 86.7 10,200 11.7
2.8 g/L 78.0 9,400 13.1
5.4 g/L 37.9 5,100 16.1
11.0 g/L l7 .0 2,500 21.0
Butanol:
Control 100 9,400 15.8
1.1 g/L 98.9 9,200 17.9
2.8 g/L 43.0 3,700 28.2
5.4 g/L 9.0 700 >30
10.3 g/L 2.0 100 >30

 

8Calculated from initial growth phase data

Optical Density (660 um)

W Density (660 n!!!)

 

 

 

  
 

 

 

 

 

 

 

08 r . r l r 44 1 I J . . r . r . r .
+C0ntrol k E

007. -
+0.1 glL / .. :

0 6 +0.25 yr. A“? ;_
' r-
+05 w . E

0.50 +1.0yr. " .-
+2.5 gli- :

0.40 E
0.3 E
0.20 . E.
- F

0.1 / .-
000. I I I l I l I I I l I I IT I h
0 24 48 72 96 120 144 168 192 216 240

Time (hrs)

Figure 11. Effect of increasing ethanol concentration during
growth of B. methylotrophicum on C0 gas.

lmIlIIIlIlllIJJIILJLII

 

0.90 +Conrrol
+01 yr.

+0.25 silo
0.70 +o.5 at.
0.60 +1” ‘4
1 +15 IIL

  
 
 
   

0.80

 

 

0.50

 

0.40
0.30
0.20

0.10
//

0.00

III I' IIII'I II I'II II 'I II I] II II II IIIII IIIIII II II II I

 

 

7‘ I I I I I I I I I I I I I
0 24 48 72 96 120 144 168 192 216 240
Time (hrs)

Figme 12. Effect of increasing butanol concentration during
growth of B. methylotrophicum on C0 gas.

87

Table 11. Balanced carbon and available electron equations for batch incubation of
stationary phase cells in the presence of alcohols during growth on C0 gas.

 

 

Alcohol
Concentration Fermentation Stoichiometrya
Ethanol:
0.0 g/L 4CO + .39AC --> 2.59C02 + .48BU + .03EH + .04BH
5.3 g/L 4CO + .38AC --> 2.62C02 + .44BU + .OSBH
Butanol:
0.0 g/L 4CO + .20AC —-> 2.48CO2 + .33BU + .06EH + .04BH + .31CM
7.1 g/L 4C0 + .30AC --> 2.52CO2 + .48BU + .14EH + .35CM

 

aSymbols used: AC = Acetate; BU = Butyrate; EH = Ethanol; BH = Butanol; CM = Dry
Cell Mass.

1.2 g/L (dry weight) for the ethanol experiment, and 5.6 g/L and 2.1 g/L for the butanol
experiment, respectively. The ethanol experiment also had initial concentrations of 0.3 g/L
ethanol and 0.9 g/L butyrate in the fermenter grown cells prior to tranferring to the CO
bottles. A cosubstrate metabolism using both CO gas and the acetate present in the broth is
evident from these calculated balances. These results indicate a potential for continued
stationary phase CO metabolism in the presence of alcohol concentrations which severely
inhibit cell growth and division. Furthermore, the data verify a mechanism in B.
methylotrophicum for exogenous acid uptake. However, as indicated, these experiments
exhibited a cosubstrate metabolism, and thus CO gas was not the sole carbon and energy
source. Another important qualiﬁcation to these stationary phase results is necesary. For
the experiment using butanol, cell production was observed concommitant with a pH drop
during the ﬁrst day, and then the optical density declined moderately over the ﬁnal 2 days
(data not shown). For the experiment using ethanol, no cell production was observed (data
not shown). This ambiguity is likely a result of the fermenter grown cells in the

experiment involving ethanol having been harvested and subsequently transferred after

88

they had already entered the stationary phase and exhibited a slight decrease in optical
density.

The effect of ethanol and butanol on the growth and performance of B.
methylotrophicum during growth on CO gas is directly relevant to any potential alcohol-
producing synthesis gas fermentation. Using continuous cell recycling, this organism has
exhibited a transient capability for production of butanol as the primary product from C0
metabolism in concentrations as high as 2.7 g/L (Grethlein et al., 1991b), a result detailed
in Chapter 4. The effects of high alcohol concentration are important if such a response is
to be optimized towards a steady-state fermentation scheme, particularly with respect to
any possible inhibition mechanism(s) on cell growth or CO metabolism. The results for
alcohol inhibition during batch cell growth clearly indicate the inhibitory effects of both
butanol and ethanol above 2 and 5 g/L, respectively, with decreases in the total CO
consumption, the total acetate and cell production, and the observed growth rates.

The precise mechanism for this effect on B. methylotrophicum was not investigated,
but a general mechanism can perhaps be inferred from published alcohol tolerance studies
with C. acetobutylicum, which are quite numerous. The major effect of alcohols,
particularly butanol, on a bacterial cell is in the cell membrane, where high butanol
concentrations have been shown to increase the membrane ﬂuidity of C. acetobutylicwn, a
phenomenon for which the cellular response is to increase the ratio of saturated to
unsaturated fatty acyl hydrocarbons (Vollherbst-Schneck et al., 1984; Baer et al., 1987).
This membrane ﬂuidization effect of butanol has several consequences, including a
disruption of the nutrient transport and energy- generation systems in the cell membrane
(Moreira et al., 1981), speciﬁc inhibition of ATPase activity (Soucaille et al., 1987), and
destabilization of the internal pH (Bowles and Ellefson, 1985). This loss of energy
generation thus results in slower growth rates and increased lag times (Bowles and
Ellefson, 1985; Moreira et al., 1981; Bajpai and Iannoti, 1988). The relative toxicity of

butanol compared to ethanol in C. acetobutylicum is approximately tenfold, a trend which

89

was observed during this study but not nearly to the same degree, and which is attributed
to the increased hydrophobicity of butanol compared to ethanol (Moreira et al., 1981).
Although the physical nature of these effects can be extrapolated to B. methylotrophicum,
it should be kept in mind that C. acetobutylicum is an alcohol-producing bacterium, with
total solvent production capacities in the 1-2% range; thus one would expect the cellular
tolerance, membrane-maintenance systems, and growth mechanisms in the presence of
butanol to be much for fully developed than for B. methylotrophicum.

Results for the stationary phase alcohol experiments are somewhat inconclusive,
particularly in light of the inoculum conditions, as desribed previously. There is not a
signiﬁcant difference in the response to ethanol between the control and the 5.0 g/L
ethanol experiment, indicating that the observed cosubstrate metabolism of acetate and CO
uptake was due not to the inﬂuence of added ethanol but to either the initial presence of
acetate or the lower starting pH of 6.4 rather than 7.0-7.1, as was used for the growth
experiments. Three variables affect the stationary phase experiments: the culture pH, the
acetate concentration, and the alcohol concentration, and distinguishing the effects of one
of these variables is not possible from this single series of experiments. The experimental
evidence does conclusively show, however, that although growth of B. methylotrophicum
is severely inhibited by increasing alcohol concentrations, the stationary phase CO
metabolism is not sensitive to concentrations up to 5.0 g/L of either ethanol or butanol.
This result is relevant to the future potential of an alcohol fermentation using B.
methylotrophicum with CO gas as the sole carbon and energy source, where a
fermentation system designed to bias the cell population towards a stationary phase
response would minimize the effects of product inhibition due to high concentrations of
ethanol and butanol.

The production of alcohols as primary products from C0 metabolism will require major
research efforts to increase the tolerance and productivity, as this study has shown

concentrations of 2—5 g/L to be inhibitory to cell growth. Since 10-15 g/L can be

90

considered a minimum concentration range required for efﬁcient alcohol separation (H. E.
Grethlein, pers. com), the overall feasibility of an alcohol fermentation directly from
C0 gas is quite low. Strain improvement and adaptation research could aid in increasing
the attractiveness of these unique biological pathways, improvements which in C.
acetobutylicum have focused on alcohol tolerance via temperature-induced membrane
stabilization (Baer et al., 1987), increases in internal stabilization factors (Soucaille et al.,
1987), and isolation of tolerant mutants by application of various mutagenic factors and

adaptation mechanisms (Hermann et al., 1985).

3.3. Effect of Acids on Growth and Product Formation

As for the alcohols, the possibility for inhibition by either acetate or butyrate exists
during a high cell density CO fermentation. However, the nature of this issue is different
for the acids, particularly acetate, in lieu of product desirability. Due to the emphasis of
this research on production of fuel and fuel enhancer compounds, alcohol production was
considered more desirable than acid production, not only because alcohols are volatile,
combustible fuels or fuel additives, but also because of separation issues. Therefore,
although acetate and butyrate inhibition of cell growth would be detrimental, feedback
inhibition of these products, particularly acetate, could be a possible mechanism for
controlling the production of alcohols during CO fermentation. If the acetate forming
pathway were inhibited by a high external acetate concentration, the cellular "pool" of
acetyl-CoA would exhibit a net increase and more acetyl-CoA would be available for
ethanol or butyryl—CoA formation. The same logic also would apply for butyrate, with a
net increase in the pool of butyryl-CoA providing more substrate for butanol production.

Another potential issue intimately involved with the acid regulation question is that of
acid uptake during growth. The optimization of B. methylotrophicum for use in a CO
fermentation has primarily depended on the regulation of CO metabolism by changes in
external pH. Although previously thought to produce only acetate from CO, decreasing the

91

fermentation pH has since been shown to regulate carbon and electron ﬂow away from
acetate and towards butyrate and butanol in extended batch fermentation (Worden et al.,
1989), continuous fermentation (Grethlein et al., 1990), and continuous fermentation with
cell recycling (See Chapter 4). Yet in all of these previous pH studies, the relative
proportioning of carbon and electrons towards the various products could only be
moderately affected, and the complete shutdown of one or more of the pathways was not
observed (Worden et al., 1989; Grethlein et al., 1990; Grethlein et al., 1991a). During
these acid tolerance studies, the overall objective was to determine if more complete
regulation could be achieved by exploiting any end-product inhibition mechanisms present
in B. methylotrophicum during growth on CO gas. Speciﬁcally, the primary objective
was to inhibit acetate formation and increase butyrate or butanol formation.

The enzymes for acetate and butyrate formation from their corresponding CoA
precursors have been examined and studied as a function of pH (Grethlein, 1989). The
two pathways consist of an acetyl (butyryl) phosphotransferase enzyme, and an acetate
(butyrate) kinase enzyme, which are known to be reversible in virro; assays involved in
measuring these enzyme activities are usually performed in reverse direction (Grethlein,
1989). With or without the inﬂuence of a mass action type phenomena, it is conceivable
that conditions can be met during CO fermentation for these enzyme pathways to operate
in the reverse direction and for the cells to take up acetate and/or butyrate and produce
either ethanol or butanol (or even butyrate, in the case of acetate uptake). It should be
emphasized, however, that such uptake would not necessarily be a reversal of the known
pathways, separate mechanisms could exist for production and consumption, as has been
suggested for C. acetobutylicum (Hartmanis et al., 1984). The beneﬁts of acid uptake
ability, if it were stable, would be numerous. Most importantly, it would allow operation a
two-stage fermentation in which acids were produced in one stage with the efﬂuent
serving as the feedstock for alcohol production in the second stage. The potential

complexity of such uptake phenomena is unclear but most likely the conditions are very

92

stressful to the cell, since ATP consumption would be required, as opposed to the normal
acid-forming routes, which produce ATP.

These questions were investigated in B. methylotrophicum for both acetate and butyrate
in batch bottle experiments similar to those described for the alcohol tolerance studies
listed above. The growth of B. methylotrophicum on CO with initial acetate and butyrate
in concentrations ranging from 0 to 12 g/L was monitored. Cell growth and product
formation were measured, and net acid production calculated. Time-course growth curves
for B. methylotrophicum cells grown in the presence of increasing acetate or butyrate
concentrations are shown in Figures 13 and 14, reSpectively. Balanced equations for these
experiments are given in Tables 12 and 13. As shown, the trends of acid uptake are
evidence for a cosubstrate fermentation mechanism in this organism for simultaneous
consumption of CO gas and either acetate or butyrate. The pH changes depicted in Figure
3 are consistent with the acetate or butyrate uptake trends shown in Table 12 ; at higher
acetate concentrations, acid uptake is observed and the media is essentially deacidiﬁed,
thus resulting in less of a pH decrease during the course of the experiment. Similarly, the
trends shown in Table 13 are consistent with the pH changes observed in Figure 14,
where at higher butyrate concentrations, a relative increase in the uptake of butyrate is
observed concommitant with an increase in acetate production, and thus a decrease in the
medium pH. These experiments are further evidence for the existence of either reversible
pathways for acetate and butyrate production or distinct production and consumption
mechanisms in B. methylotrophicum, with the formation and uptake reactions both
exhibiting a strong dependence on the corresponding external acid concentration.

To test the effect of increased osmotic stress due to the experimental set-up, where to
attain a constant initial pH of 7.0-7.2, the bottles containing acetic and butyric acid were
pH-adjusted with N aOH solution, the effects of sodium on the growth of B.
methylotrophicum were investigated. The molar addition of sodium was a maximum at an

acetate concentration of 12.3 g/L, or 0.205 M, which corresponds to an addition of

93

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Table 12. Balanced carbon and available electron equations for batch incubation in the
presence of acetate during growth on CO gas.

 

 

Acetate Concentration.
Fermentation Stoichiometrya
0.0 g/L 4C0 --> 2.08CO2 + .61AC + .043U + .OIIBU + .03EH + .50CM
2.3 g/L 4CO --> 2.21CO2 + .30AC + .14BU + .OZIBU + .02EH + .52CM
4.5 g/L 4CO --> 2.20C02 + .29AC + .15BU + .OlIBU + .01EH + .55CM
8.4 g/L 4CO + .11AC --> 2.38CO2 + .32BU + .021BU + .03EH + .45CM
12.3 g/L 4CO + .18AC --> 2.41CO2 + .34BU + .021BU + .03EH + .44CM

 

aSymbols used: AC = Acetate; BU = Butyrate; IBU = Isobutyrate; EH = Ethanol;
CM = Dry Cell Mass.

Table 13. Balanced carbon and available electron equations for batch incubation in the
presence of butyrate during growth on C0 gas.

 

Butyrate Concentration.
(g/L) Fermentation Stoichiometrya

 

0.0 g/L 4C0 --> 2.10C02 + .57AC + .063U + .OlIBU + .01EH + .48CM
2.2 g/L 4C0 + .O7BU --> 2.01C02 + .85AC + .OlIBU + .02EH + .02BH + .42CM
4.7 g/L 4C0 + .28BU --> 1.84C02 + 1.25AC + .OZIBU + .02EH + .03BH + .57CM
8.7 g/L 4C0 + .34BU --> 1.77C02 + 1.44AC + .OIIBU + .02EH + .03BH + .52CM
12.5 g/L 4C0 + .42BU --> 1.69C02 + 1.68AC + .OlIBU + .02EH + .03BH + .44CM

 

aSymbols used: AC = Acetate; BU = Butyrate; IBU = Isobutyrate; EH = Ethanol;
BH = Butanol; CM = Dry Cell Mass.

96

0.0102 acid equivalents. Thus, to neutralize this acid, a minimum of 0.012 base
equivalents was required, in the form of NaOH. In 50 mL liquid, this is equivalent to 0.24
mol Na/L, or, in the form of NaCl, approximately 0.25 mol NaCl/L. An experiment
testing the effect of NaCl was thus conducted over the concenteration range of 0-1.0 g/L,
and the effect on the growth on B. methylotrophicum determined during CO metabolism in
pressure tubes. Results from this experiment are shown in Figure 15, where only at a
NaCl concentration of 1.0 g/L is signiﬁcant inhibition of cell growth observed, mostly in
the form of an increased lag phase, a result also seen at 0.75 g/L NaCl. Therefore, from
these results, the effect of osmotic stress from Na addition was determined not to inﬂuence
cell growth during the tolerance studies with acetate and butyrate.

A metabolic precedent for the phenomenon of acid uptake can be found in E. Iimosum
grown on methanol, where the production of acetate can be decreased by adding acetate to
the growth medium, with higher concentrations reversing the reaction and causing acetate
uptake (Pacaud et al., 1986a; Pacaud et al., 1986b). Fin-thermore, the same effect has been
observed with butyrate, with increasing butyrate concentrations causing an increase in the
relative production of acetate (Loubiere et al., 1986), a trend observed here for butyrate
uptake by B. methylotrophicum. Similar phenomena have also been observed for
Closm'dium tyrobutyricum grown on glucose (Michel-Savin et al., 1990). The authors of
these studies state that the observed acid uptake occured by reversal of the acid-forming
pathways, although no enzyme activity studies were conducted.

The uptake of external acetate and butyrate during batch growth of B .
methylotrophicum raises the important question of the potential for a steady-state uptake
response over long time periods, since this mechanism would place additional stress on
the cellular energetics. Assuming the uptake of external acids proceeds via a reversal of the
known production pathways, the uptake of acetate, on a basis of acetyl-CoA formation,
requires 1 ATP for every 1 acetate consumed or 1/2 an ATP for every 1 butyrate

consumed. With this stoichiometry, one would predict that acetate production from

97

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98

butyrate uptake would be more favorable than butyrate production from acetate uptake, if
only from an ATP formation standpoint (based on SLP only). This trend is consistent with
the experimental data, which indicate a higher afﬁnity for butyrate uptake, observed at 2.2
g/L and above, than for acetate uptake, observed at 8.4 g/L and above. This hypothesis
may also explain why a major increase in butanol production is not observed during
butyrate uptake, since the butanol pathway from butyrate yields no ATP, but conversion
of butyrate to acetate does. The importance of energetics during CO metabolism can be
seen in the metabolic pathways for l-carbon metabolism (Zeikus et al, 1985) as well as
acid and alcohol formation from C0 (Grethlein et al., 1991b) , which have been

previously elucidated for B. methylotrophicum. The ATP consuming reaction in the CO

pathway is the formation of formyl-H4folate, and thus ultimately the methyl group of
acetyl-CoA, which is catalyzed by formyl-H4folate synthetase (Kerby and Zeikus,
1987b). The ATP producing steps are the kinase reactions for formation of acetate and
butyrate. From these known pathways, the molar ATP and electron yields arising from
substrate-level CO metabolism can be calculated, as shown in Table 14. A critical

Table 14. Theoretical molar ATP and electron yieldsa from substrate-level
phosphorylation for products derived during CO metabolism in

 

 

Butyribacterium methylotrophicwn.
Product ATP Yield Electron Yield
Acetate O O
Butyrate -1 -2
Ethanol -l -2
Butanol . -2 -4

 

aMoles ATP and Electrons per Mole of Product

99

observation arising from these calculations is that only acetate production requires no net
input of energy or electrons; all other products from C0 metabolism consume ATP and
electrons. This observation is strong indirect evidence for the existence of an electron
transport phosphorylation mechanism in B. methylotrophicum, since products other than
acetate, including cell mass, are observed from C0 metabolism. Preliminary evidence has
been obtained to support the existence of such an ATP generating mechanism based on
ETP (Zeikus et al., 1985), but conclusive experimental evidence has not yet been
obtained. The required electrons for production of butyrate, alcohols, and cell mass from
C0 presumably arise from the oxidation of C0 to C02, which would explain the increase
in C02 yields observed experimentally (Grethlein et al., 1990) as CO metabolism is
regulated away from acetate production and towards butyrate (or alcohol) production.

The growth experiments to determine the effect of external acid concentration on B.
methylotrophicum during CO metabolism have provided a possible mechanism to regulate
carbon and electron ﬂow away from acetate production. A capacity for both acetate and
butyrate uptake was observed as a function of increasing acid concentration during growth
of the organism on C0 gas. This observed acid uptake thus yielded a cosubstrate
fermentation proﬁle, with either CO and acetate or CO and butyrate serving as the carbon
and energy sources. The uptake of acids was furthermore observed concommitant with
CO consumption, and thus it is reasonable to assume that acetate and butyrate serve as
precursors for acetyl-CoA and butyryl-CoA formation, but not as energy sources for
oxidation of acetyl-CoA to C02, since this would reverse the carbon and electron ﬂow
observed in situ during CO metabolism (Kerby et al., 1983).

This study has further shown the metabolic versatility of B. methylotrophicum with the
observed ability for acetate and butyrate cosubstrate fermentations with CO gas. These
acid uptake mechanisms could have significant impact on the design of a butyric acid
fementation using CO gas, with a ﬁrst stage fermentation producing acetic acid from C0,

and a second stage fermentation producing butyric acid from C0 and acetic acid. Such a

100

scheme could be more favorable than a direct CO to butyric acid fermentation, where

complete termination of acetic acid formation has proven to be extremely difﬁcult.

3.4. Effect of Sulﬁde on Growth and Product Formation

An attractive advantage of biological synthesis gas conversion is that intensive sulfur
gas removal and synthesis gas upgrading may not be required. With sulfur-tolerant
microorganisms such as B. methylotrophicum, the 0.5-2 volume % sulfur gases present in

a typical coal-derived synthesis gas will potentially pass through the process as inert
components. Because the CO (and H2 in the MBI Process) will be removed from the CO

gas stream, a reduced volume of gas having a higher H28 concentration will exit the
process. Sulfur removal costs will thus be reduced, because the mass transfer driving
force will be greater, while the volume to be processed will be less. Furthermore, B.
methylotrophicum does not require a strict CO/Hz ratio, and thus gas shift operations
would potentially be unnecessary.

Although anaerobic media requirements provide a high level of sulfate and sulﬁde-
containing chemical species, the specific effects, if any, of H28 gas on B .
methylotrophicum during growth and product formation are unknown. Therefore the
effects of H28 gas was determined during batch growth on CO, with initial mole fractions
of H28 between 0.005 and 0.02--a range typical of actual compositions in coal-derived
synthesis gas mixtures (V erma, 1978a; McCullough et al., 1984). In order to quantify the
sulfur-tolerance of B. methylotrophicum, the growth and product-formation properties
during CO metabolism were determined in the presence of sulﬁde. These results are
necessary to evaluate the potential economic beneﬁts of the synthesis gas bioconversion
process.

To determine the eﬁ’ects of sulﬁde in an experimental biological system, the equilibrium
partitioning and dissociation chemistry of sulﬁde species must be addressed. The liquid-
phase concentration of H28 in equilibrium with a given H28 partial pressure (szS) can

101

be estimated accurately for pressures less than two atmospheres using Henry's Law
(Carroll, 1990):
[H28] = (KH)(PH23)-
Values for the Henry's law constant (KH) published in the International Critical Tables
and Perry's Chemical Engineer's Handbook were found to be incorrect by several orders
of magnitude. However, correct values were estimated using a published correlation
(Carroll, 1990) and were used for calculations made during these experiments.
In aqueous solution, H28 acts as a diprotic acid, as shown below,

H28 <=> H+ + HS' pKa1 = 7.02

as <===> H+ + sZ- pKa2 = 13.9
where the ratios of the various sulﬁde species may be calculated using the Henderson-
Hasselbach equation:

pH=pKa1+logtms-1/mzs]) pH=pKa2 +log([82-1/[Hs-]).

Because B. methylotrophicum operates in a pH range of 5.5 to 7.2, the concentration of
82" will be negligibly small. Based on these relationships, the following equation was
developed to give C, the total equilibrium sulﬁde concentration (including all species) in
the liquid phase following an initial loading of n moles of sulﬁde:

n[1 + 10(pH-pKa )1
c =

 

VG/KHRT + an + MPH-pm]
where VG and VL are the volumes of the gas and liquid phases, respectively. The
absolute temperature is represented as T, and the gas constant as R. This equation is valid
regardless of the form of sulﬁde added (e.g., H28 or Na28). At the low pressures
involved in these experiments, the number of moles of H28 added can be estimated using
the ideal gas law with the appropriate volume of gas added. The liquid-phase

concentrations of individual species may also be calculated, as shown below:
[1123] = C/[l + 10(pH-pKa )1

102

HS' = C - [I-IzS]
pHZS = [HZS]/KH.

The equations used to calculate sulﬁde concentrations in the gas and liquid phases are
based on the assumption of phase equilibrium. To determine whether this assumption was
reasonable during the sulfur-tolerance studies, a batch experiment was conducted to
estimate the time required for equilibrium to occur after the addition of H28 to gas
headspaces in bottles containing distilled water. This experiment was performed in
pressure bottles, which were well mixed by shaking after addition of H28 gas. As the
H28 partitions into the liquid phase and dissociates, the pH of the distilled water would be
expected to decrease. Equilibrium was assumed to have occured when the pH stopped
decreasing. The results for this experiment are shown in Table 15, for duplicate bottles
initially containing 2%, 5%, and 10% H28 gas in the headspace (by volume). For all of
the concentrations tested, the pH ceased decreasing after 15 minutes, with over 90% of
thepH change occuring during the ﬁrst 2 minutes.This is a rapid change relative to the

120-160 hr duration of the growth experiments with the sealed bottle systems. Therefore

Table 15. Experimental pH proﬁles showing dynamics of H28 dissolution.

 

 

Trme pH of Sample
(min) 2% H28 5%1-128 10% H25
0 6.43 6.46 6.51
2 5.52 5.42 5.24
15 5.50 5.41 5.20
30 5.56 5.43 5.19
45 5.65 5.43 5.26
60 5.71 5.54 5.42
90 5.76 5.55 5.44
120 5.80 ND 5.44

 

ND denotes "not determined"

103

the assumption of phase equilibrium was deemed valid for the sulfur tolerance
experiments.

The sulfur tolerance of B. methylotrophicum grown on C0 gas was investigated from
the standpoint of total liquid-phase sulﬁde concentration. This organism requires a media
reducing agent to maintain anaerobic conditions, and in the phosphate-buffered media,
Na28 is typically used during batch culture (Lynd et al., 1982). The standard amount of
sulfide initially added to the liquid is 0.083 g/L, which corresponds to an equilibrium
liquid phase sulﬁde concentration of 1.7 mmol/L. This value is an amount equivalent to
contacting a 102 mL 3.2% H28 gas phase with a non-reduced liquid media volume of 50
mL, which are the standard experimental conditions. The resulting equilibrium H25 gas
phase concentration arising from the control addition would be 1.13%. Thus, this
organism is inherently tolerant to sulﬁde species.

Using an initial liquid-phase sulﬁde concentration of 1.7 mmol/L as a control value,
three distinct sulﬁde tolerance experiments were conducted: the ﬁrst using the control
amount of Nazs plus addition of increasing amounts of H28 gas, the second using
increased amounts of Na28 relative to the control, and the third using decreased amounts
of Nazs relative to the control. These experiments were all conducted at an initial pH of
7.0. The equilibrium liquid phase sulﬁde concentration at this initial pH can be considered
to be a maximum, since during the course of the CO fermentation the pH decreases, and
thus promotes formation of H28 gas as determined by the dissociation chemistry. Figures
16 and 17 show the growth and acetate production time-course proﬁles for batch growth
of B. methylotrophicum during the ﬁrst experiment. The values of percent H28 gas
indicate initial equilibrium gas phase concentrations. The additional sulﬁde severely
inhibited both cell growth and acetate production in the 3.2%, 5.2%, and 8.0% runs.
Molar and speciﬁc yield values for acetate production during this experiment are shown in
Figure 18. As the initial equilibrium sulﬁde concentrations increased, and the resulting cell

 

 

 

 

 

 

 

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Time (hrs)

Figure 16. Effect of initial 1-128 gas phase concentration during
growth of B. methylotrophicum on C0 gas.

 

 

 

 

 

 

 

 

 

250 l I I l 1 I 1 l J I I I 1 v _
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. +2.16% :
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0.0 — I I I ‘l T I 1 I I
0 24 48 ‘72 96 120 144 168 192

Time (hrs)

Figure 17. Effect of increasing H28 concentration on acetate
production during growth of B. methylotrophicum.

105

Molar Yreld (mol Acetate/mol CO)

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106

growth rates decreased, the speciﬁc and molar yields for acetate increased by
approximately 50%.

For the second experiments, the time-course profiles for cell density are shown in
Figures 19 and 20. The initial liquid sulﬁde concentrations shown in these ﬁgmes are pre-
equilibrium values, calculated assuming no equilibrium partitioning. At concentrations
below the control (0.083 g/L) value, shown in Figure 19, there was no signiﬁcant effect
on cell growth, except that the lowest concentration tested did show a slight increase in the
growth rate. At concentrations above the control, shown in Figure 20, the growth rate and
corresponding cell densities decreased with increasing sulﬁde concentration. The data for
acetate, and minor butyrate, formation during these experiments paralleled the trends for
cell density and cell growth and are therefore not shown. However, the speciﬁc and molar
yields of acetate and butyrate were signiﬁcantly inﬂuenced by the concentration of sulﬁde,
and are shown in Figures 21 and 22. The two experiments are shown in separate ﬁgures
because of the separate control runs for each experiment. As the sulﬁde concentration was
increased, the acetate yields initially decreased, reached a minimum, and then increased.
The butyrate yields exhibited the opposite trend.

A summary of the results which used Nazs as the sulﬁde source is shown in Table 16,
which relates the total initial, equilibrium, liquid-phase sulﬁde concentration to the relative
cell growth and corresponding gas phase H2S concentrations. Initial equilibrium gas
phase concentrations approaching 1.4% did not signiﬁcantly inhibit cell growth. Reduced
rates of growth were observed between 1.4% and 3.3%, and concentrations above 3.3%
were toxic to B. methylotrophicum.

The experimental results for sulﬁde tolerance obtained with both H28 gas and Na28
serving as the sulfur sources have determined a practical concentration limit for liquid
phase sulﬁde species during cell growth and metabolism. The growth results for B.
methylotrophicum were based upon the initial liquid concentration of sulﬁde species that is
typically used as a reducing agent in the phosphate-buffered media (1.7 mmol/L). The

2.0 '
1.8
1.6
1.4
1.2
1.0

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107

 

 

 

 

 

 

 

V l I I j

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24 48 72 96 120
Time (hrs)

Figure 19. Effect of initial sulﬁde concentration during growth

 

 

of B. methylotrophicum on C0 gas.

 

 

 

 

 

 

 

1.6 “#4 ‘ l ‘ I—
3 +983 w 5
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Figure 20. Effect of initial sulﬁde concentration during growth

of B. methylotrophicum on C0 gas.

Speciﬁc Yield (mol Product/mol Cell)

108

 

 

 

 
 

 

 

 

 

 

 

 

 

 

   
 

 

 

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0.00 0.02 0.04 0.06 0.08 0.10
Initial Liquid Sulﬁde Concentration (g/L)
Figure 21. Speciﬁc and molar yields as a function of initial
sulﬁde concentration dming growth of B.
methylotrophicum on C0 gas.
1.30 .. I I I I I I I I I I I I I I I I I I I I I I I I 0.16
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0.00

 

 

0.06 0.12 0.18 0.24 0.30

Initial Liquid Sulﬁde Comentr'ation (g/L)

Figure 22. Speciﬁc and molar yields for product formation during

growth of B. methylotrophicum on C0 gas.

109

Table 16. Effect of liquid phase sulﬁde concentration on cell growth of B.
methylotrophicum grown on C0 gas at an initial pH of 7.0.

 

 

Total Sulﬁde Species Equilibrium Gas % Maximum OD
in Liquid Phase Phase H28 Concentration Obtained

_(mellLl (mQL%) (% Com—
0.479 0.32% ~102%
0.890 0.59% ~100%
1.301 0.86% ~100%
1.71 1 1.13% 100%
2.053 1.36% ~93%
3.012 1.99% ~72%
4.107 2.71% ~48%
5.203 3.32% ~27%

 

experiment with excess addition of H28 gas to the control media showed severe inhibition
of cell growth in the presence of greater than 2% H25 gas. The speciﬁc mechanism of this
inhibition is not clear, but an effect on the molar and speciﬁc yields is apparent from the
data presented in Figure 18, where high sulﬁde concentrations appear to favor acetate
formation. This result is most likely a cellular response to the toxic effects of high sulﬁde
concentrations, a response which may require more metabolic energy and thus would
favor acetate production at the expense of cell mass production. Whatever the speciﬁc
mechanism, it is evident that the overall effect of sulﬁde is dependent on the physiology of
the bacteria; studies with P. productus and R. rubrum have shown almost an order of
magnitude greater sulﬁde tolerance, although under somewhat different experimental
conditions (Vega et al., 1990).

The data obtained using N a28 as the sulfur source show that the initial, pre-equilibrium

control concentration of 0.083 g/L total sulﬁde is on the border between tolerance and

110

inhibition, with lower concentrations having little or no effect on cell growth (Figure 19),
and with higher concentrations leading to inhibition of cell growth (Figure 20). The
speciﬁc and molar yields of acetate and butyrate, plotted in Figures 21 and 22, are affected
by sulﬁde concentration in a complicated manner. At low sulﬁde concentrations, the
acetate yields are higher, decrease to a minimum at an intermediate concentrations, and
then increase again as the sulﬁde concentration is increased further. The butyrate speciﬁc
and molar yields exhibit the opposite trend. Currently these trends cannot be adequately
explained.

It has been previously shown that Na2S addition to C. acetobutylicum cultures results
in a metabolic shift resulting in enhanced butyrate production (Rao and Mutharasan,
1988). The addition of Na28 to Thermoanaerobacter ethanolicus cultures has also been
shown to inhibit cell growth at high concentrations (Rao et al., 1987). These ﬁndings are
consistent with our ﬁndings for B. methylotrophicum, where the liquid phase sulﬁde
concentrations above 3.0 mmol/L signiﬁcantly reduce cell growth and product formation.
However, the effects of sulﬁde on stationary phase CO metabolism could potentially be
much less severe, as was observed for high alcohol concentrations which inhibited growth
during the exponential phase but allowed CO metabolism to proceed during the stationary
phase. Regardless of this speculation, since the equilibrium partial pressures of H28 gas
above these liquid concentrations are in the range of 1-2%, with proper process design the
influence of H28 on the growth and metabolism of B. methylotrophicum should not
create any problems during coal-derived synthesis gas conversion, and should
substantially reduce the requirements for stringent sulfur gas removal prior to the

fermentation, relative to the typical chemical process requirements.

3.5. Media Development Studies
To apply cell recycling technology to a continuous fermenter for increasing cell

densities and volumetric productivities during C0 fermentation, as detailed in Chapter 4, a

111

modiﬁed phosphate-buffered medium was needed to ensure that during high cell density
culture, CO gas-liquid mass transfer is still the rate limiting factor. Without such
modiﬁcations, nutrient limitations could drastically alter the observed kinetics and make
comparison with previously obtained chemostat data extremely difﬁcult. However, the
potential enhancement must not interfere with the premise that CO is the sole carbon and
energy source during B. methylotrophicum metabolism. A media development effort was
therefore conducted in order to increase the amount of nutrients and minerals in the
fermentation media, without adding alternate sources of carbon or energy.

The original phosphate-buffered basal medium used for culturing B. methylotrophicum
was detailed in Chapter 2. Carbon, in the form of CO, was chosen as the limiting nutrient.
For B. methylotrophicum, the H2 is obtained from H20, and phosphorous will not be
limiting in a phosphate-buffered environment. Approximate cellular elemental
compositions of nitrogen (14% w/w) and sulfur (1% w/w) were obtained from available
E. coli data (Bailey and Ollis, 1986). Total required concentrations of these elements in a 1
OD broth solution can then be calculated by multiplying by the dry weight concentration of
0.334g/L-OD (Lynd et al., 1982), yielding 0.00334 g/L sulfur and 0.04816 g/L nitrogen
required. The original phosphate-buffered medium contains .0156 g/L and .252 g/L of
these elements, respectively. Thus, assuming a 10-fold increase obtained in cell density,
the required amounts of these elements would be 0.0334 g.L sulfur and 0.4816 g/L
nitrogen. Therefore, as a minimum, the concentrations of these two elements should be
doubled in the enhanced medium. From the available chemicals, the concentration of
sulfur was approximately tripled using ammonium sulfate, and the concentration of
nitrogen was approximately doubled using ammonium sulfate and ammonium phosphate,
with the ammonium phopshate added to the phosphate buffer as a partial substitute for
dibasic potassium phosphate. Abitrarily, the concentration of vitamins was increased 10-
fold, and the concentration of trace minerals was increased 5-fold. A 2-fold increase in the

yeast extract concentration was also included, as determined below.

112

Experiments to determine the effect of these modiﬁcations on the growth of B.
methylotrophicum were conducted in batch bottles with a control of the original medium,
experimentals for each of the separate additions, and with all of the additions together. The
time-course growth curves are shown in Figure 23 for the various medium supplements,
and in Figure 24 for the tests with increasing concentrations of yeast extract. The cell
growth rates in the presence of all the supplements shown in Figure 23 are similar, and
only the 2-fold nitrogen test indicated a signiﬁcant increase in the ﬁnal cell density relative
to the control. The lag phase on some of the additions are slightly increased. It was thus
determined that these additions do not signiﬁcantly effect the typical growth response of
B. methylotrophicum when grown on C0 gas. The effect of increasing yeast extract,
however, was signiﬁcant, as shown in Figure 24, particularly for the 3-fold and 5-fold
increases, which exhibited reduced lag phases, slightly increased growth rates, and higher
cell densities. The 2-fold increase exhibited a similar growth rate and ﬁnal OD relative to
the control, but with a reduced lag time. It was thus decided to incorporate a 2-fold
increase in the yeast extract concentration, from 0.1% (w/v) to 0.2% (w/v). This increase,
particularly in view of anticipated 10-fold increases in cell density during cell recycle
fermentations, will result in a net decrease in the mass ratio of yeast extract to cells, and
thus cannot be argued to support cell growth alone, an argument which was refuted

previously for cells grown on C0 gas and 0.05% yeast extract (Lynd et al., 1982).

Optical density (660 nm)

Opyical density (660 mu)

 

  

 

  
   
    
   
    

 

 

    

 

      

 
      

  

 

1 20 I I I I I I I I I I I 1 I I
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+All Additions :

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l
48 72 96 120 144 168
Time (hrs)

I
0 24

Figure 23. Effect of media supplements on the growth of B.

 

 

   
    

 

 

 

 

 

 

methylotrophicum on C0 gas.
1.40 I I I J I L I I I L I I
+larln> E
1.20 -
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0.60 E
0.40 -
0.20 5'
0.00 h ' I I I I I I I I
0 24 48 72 96 120 144
Time (hrs)

Figure 24. Effect of increasin yeast extract concentration on the
growth of B. me ylotrophicum on C0 gas.

CHAPTER 4

CELL-RECYCLE FERMENTATION STUDIES

4.1. Objectives

Biological processing of synthesis gas requires the application of microorganisms
capable of both efﬁcient metabolism and sustained growth on CO, with the objective of
producing organic acids or alcohols. Research over the last few years in several
laboratories has identiﬁed the few critical problem areas which must be addressed and
overcome for the ultimate success of a bioconversion concept. These areas include low
rates for gas-liquid mass transfer of both CO and H2, low speciﬁc rates of growth, and
also a potentially wide range of multiple products resulting from C0 metabolism. Another
disadvantage with biological systems is the difﬁculty associated with maintenance of
microbial populations in desired metabolic states, particularly if these states are not optimal
for cell growth and viability. These issues all affect fermenter productivity, which has
been consistently shown to be much too low for an economically viable process. The
major objective of the cell recycling studies contained in this chapter was to address these
concerns.

As discussed below, the previous continuous culture results suggested that to maximize
butyrate production, long-term operation of the acidogenic CO fermentation at low pH
values is required. These key issues were addressed by incorporation of both gas and cell
recycle systems into the chemostat CO fermentation system. Such a system, when used
during continuous CO fermentation, potentially allows operation at low pH, attains much
higher cell densities than with conventional chemostat systems, and thus can achieve large
increases in fermenter productivity.

Extended batch fermentations with B. methylotrophicum determined that fermentation
pH has a signiﬁcant inﬂuence on the relative proportion of acetate and butyrate produced

114

115

from CO or CH30H and CO (Worden et al., 1989; Grethlein, 1989). Continuous, steady-
state CO fermentations at pH values between 5.5 and 6.8 also showed this trend, with
butanol and ethanol production occuring at lower pH values (Grethlein, 1990). These
continuous results, showing butanol and ethanol formation directly from C0, identiﬁed a
potential mechanism for direct bioconversion of synthesis gas to alcohols. Carbon and
electron balance calculations, based on the broth product composition as a function of
fermentation pH during steady-state operation, yielded several signiﬁcant trends for these
experiments. As shown in Tables 17 and 18, there was a trend towards more reduced
product formation, speciﬁcally butyrate and butanol, as the fermentation pH was reduced.
Also evident was a drop in the cell growth as the pH was reduced to a value of 6.0 with
further reduction observed at a pH of 5.5 (washout was observed at a pH of 5.0). These
trends deﬁne the major experimental obstacles for fermentation development of this
biological system: low fermentation pH induces butyrate and alcohol formation but inhibits
cell growth and thus reduces fermenter productivity.

The low speciﬁc growth rate of this organism, as indicated by a 12 hour doubling time,
required a very low dilution rate for steady-state operation during previous continuous
culture. Thus, although this operation was achieved, volumetric productivities for acids
and alcohols were very low, ranging approximately between 0.00] g/L hr for the alcohols
and 0.015 g/L hr for the acids. Production of acids and alcohols was achieved at
concentrations of approximately 0.05% to 0.1% (w/v) for the acids and 0.005% to 0.01%
(w/v) for the alcohols (Grethlein et al., 1990), concentration levels which are of scientiﬁc
interest only. Higher dilution rates and cell densities are therefore required in order to
increase product concentrations and volumetric productivities.

This dissertation research was focused on fermentation development in order to achieve
a stable, long-term, and highly productive CO fermentation with B. methylotrophicum. As
desribed above, the major method used to achieve this objective was the application of cell

recycle membrane technology to a conventional continuous fermentation system. This

116

Table 17. Continuous, steady-state CO fermentation stoichiometries as a function of pH

for B. methylotrophicuma.

 

 

pH Fermentation Stoichiometry
6.8 4co ---> 2.09C02 + .63CH3COOH + .043C3H7COOH + .027C2H50H + .43Cells
6.5 4C0 --> 2.13C02 + .56CH3COOH + .082C3H7COOH + .026C2H50H + .37Cells
6.0 4co ---> 2.27c02 + .30CH3COOH + .161C3H7COOH + .032C2H50H

+ .029C4H90H + .31Cells
5.5 4C0 ---> 2.18C02 + .40CH3COOH + .154C3H7COOH + .40Cells

 

aFrom Grethlein et al., 1990.

Table 18. Continuous, steady-state CO fermentation product concentrations as a function

of pH for B. methylotrophicuma.

 

 

Product Concentrations (g/L)
pH CH3COOH C3H7COOH C2H50H C4H9OH Cells
6.8 0.860 0.086 0.028 ND 0.248
6.5 1.055 0.227 0.037 Trace 0.284
6.0 0.689 0.536 0.056 0.081 0.286
5.5 0.475 0.266 Trace Trace 0. 196

 

aFrom Grethlein et al., 1990. Trace indicates amounts < 0.5 mmol/L, ND = Not Detected.

117

technology allows high cell densities by restricting cell washout during continuous
operation, while allowing soluble fermentation products to leave the system. In general,
the application of cell recycling allowed operation at a wider range of physical and
environmental conditions without any washout restrictions, compared to conventional

chemostat culture.

4.2. System Set-up and Design

To increase the overall productivity of the CO fermentation at low pH values, and to
bias the cell population towards a stationary (resting) state metabolism, a cell recycling
system was integrated with a continuous fermentor, as shown in Figure 25. This system
was designed for higher CO gas mass transfer by using a pump to continuously remove
the headspace gas and sparge it into the fermentor broth. CO gas was provided to the
fermentation continuously. An enhanced phosphate buffered media, modiﬁed by
increasing the vitamin, mineral, and nutrient contents as described in Chapter 3, was used
to support higher cell densities while still maintaining CO as the limiting nutrient, as in all
previous studies. Experiments were conducted at a wide range of pH values, in order to
achieve high cell density, steady-state CO fermentations. Steady-state was deﬁned as
liquid product concentration deviations of not more than 15% from the mean value over a
period of 2-3 residence times, or approximately 6-9 days.

In order to adequately supply the fermentation with enough CO to support the expected
increases in cell density, high mass transfer rates for CO were obtained by using a high
impeller speed and with a CO gas recycle pump. These developments allowed high gas
ﬂows into the liquid and provided increased dispersion of the dissolved gas. Furthermore,
macroporous frits were used to minimize bubble size for the sparged CO gas.

The major investigation using the cell recycle fermentation system described above was
the effect of fermentation pH. This variable had already been shown to be highly

inﬂuential on production rates and product speciﬁcity during CO metabolism. Cell

118

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recycling allowed investigation at pH values lower than those obtainable during
conventional continuous culture, where reductions in the speciﬁc growth rate of B.
methylotrophicum at pH values below 6.0 were previously observed (Grethlein et al.,
1990). Fermentations were conducted at a range of pH values from 5.0 to 6.8 and the
fermentation proﬁle measured, including product formation and cell density. Emphasis
during the cell recycle experiments was on achieving steady-state response at each pH
value tested. The dilution rate for these studies was 0.015 hrl, which is the same as for
the previous chemostat experiments, in order to allow direct comparison. All of the
fermentations were operated at a conversion rate of approximately 20%, with
approximately 5-7% deviation ﬁom this value observed. Product formation data allowed
calculation of overall stoichiometries, yield coefﬁcients, volumetric productivities, and
other kinetic parameters.

A major limitation with the experimental system design was with the reliance on the
carbon and electron balances for calculation of cell production and CO consumption. The
use of these balances assumes 100% conservation of carbon and electrons calculated from
the fermentation data. Thus, since all of the experimental data is ultimately normalized to
the CO consumption, any error in the system or in the analytical methods will ultimately be
reﬂected in the calculated values for cell mass production. These limitations were
somewhat voluntarily accepted; the methods for measuring direct CO consumption (gas
phase tie-component analysis) and for measuring direct cell production (viability plate
assays) are very labor intensive and were deemed impractical dlning this study.

Another important point concerning this experimental system is the nature of its
development and operation over long-term time periods. With a 12 hour doubling time,
each fermentation conducted required approximately 3 weeks just to raise the cell density
high enough before adjusting the controlled variable, whether it was fermentation pH or
the addition of a cosubstrate. Achieving steady-state was then a matter of precise

monitoring and control (and also luck) since small ﬂuctuations in operating conditions

120

over days of operation could cause signiﬁcant perturbations in the system. The end result
was that for each fermentation, an average of 6-8 weeks was required either to achieve
steady-state, or to obtain a complete picture of the fermentation response. Since only 1 and
sometimes 2 experiments were conducted at a time, there was some amount of operating
expertise gained over the course of this study, which is generally reﬂected in the cell
densities for the fermentations (the later experiments exhibited higher cell densities and
generally greater stability). These limitations and experimental characteristics must
therefore be taken into account when comparing the results from the various cell recycle
fermentations. In general, the analysis of experimental results in this chapter reﬂects this
point, and relies more heavily on fermentation stoichiometries than on direct comparison
of product concentrations.

One aspect of typical cell recycle systems which was not a factor during these studies
was the problem of membrane fouling or life. The mineral salts nature of the phosphate
medium and the lack of extracellular slime or polymer production by B. methylotrophicwn
contributed to "clean" fermentation broths which did not have a great afﬁnity for the
cellulosic membranes. Even in systems which were operated for up to 6-8 weeks,
backﬂushing and membrane cleaning operations were not required. Of course, the lack of
fouling was largely due to a very low ﬂux through the membrane, since the liquid

residence times of the CO fermentations were approximately 67 hr.

4.3. Results for Continuous Cell Recycle CO Fermentations

The cell-recycle system was operated between fermentation pH values of 5.1 and 7.2.
Three distinct metabolic regimes were observed, as presented in Table 19. At pH values of
7.2, 6.8, 6.4, and 6.0, steady-state was achieved. At pH values of 5.75 and 5.5,
oscillations in the concentrations of several fermentation products were observed. A viable
culture could not be maintained during Operation at pH values below 5.5. The results for

each of these regimes are discussed in detail below.

1 2 1
Table 19. Overall culture response as a function of fermentation pH during continuous

 

 

cell recycle fermentation.

Emantan'onnﬂ War:
7.2 Steady-State
6.8 Steady-State
6.4 Steady-State
6.0 Steady-State
5.75 Oscillations
5.5 Oscillations
5.35 Death
5. 1 Death

Wank:

The steady-state CO fermentations exhibited trends in product composition similar to
previous experiments without cell recycling (Grethlein et al., 1990). Table 20 shows the
steady-state fermentation stoichiometries for these experiments. The coefﬁcients were
determined using carbon and available electron balances. A consistent trend of decreasing
acetate production and increasing butyrate and alcohol production is evident as the pH
decreases from 7.2 towards 6.0. At a pH of 6.0, a shift from acid to alcohol production is
indicated, since the proportional amounts of both acetate and butyrate decrease while the
relative alcohol production increases. By dividing the coefﬁcients for the products by the
coefficient for CO in the fermentation stoichiometries shown in Table 20, the molar yield
coefﬁcients can be obtained. These are shown in Table 21. The relative production for
each compound can be seen as a function of decreasing pH, with decreasing acetate and
cell yields, increasing butyrate yields, and at a pH of 6.0, and a 40-50% increase in the
alcohol yields, presumably at the expense of the corresponding acids. These trends are

further exempliﬁed by the partitioning of carbon and energy (electrons) available in the CO

122

Table 20. Steady-state fermentation stoichiometries as a function of pH for cell recycle

fermentation of CO gas.

 

6.8

6.4

6.0

E 'S'l‘

4co -..> 221c02 + .410CH3COOH + .rosc3rl7coorr + .019C2H50H + .032c4ngorl + 387CELLS
4C0 -..> 2.25c02 + 334cn3coon + .124C3H7COOH + .025C2H501-l + .040C4H90H + 377caus
4co -..> 226C02 + 316CH3C00H + .152c3rl7coorr + .018C2H501-l + .032C4H90H + 279031.15

4C0 --> 232C02 + .260CH3COOH + .142C3H7COOH + .050C2H50H + .055C4H90H + .279CELLS

 

Table 21. Continuous, steady-state CO fermentation molar yield coefﬁcents as a function

of pH for B. methylotrophicum.

 

 

Molar Yield Coefﬁcients
pH C02 CH3COOH C3H7COOH C2H50H C4H9OH Cells
7.2 0.552 0.102 0.026 0.005 0.008 0.097
6.8 0.562 0.084 0.031 0.006 0.010 0.094
6.4 0.565 0.079 0.038 0.005 0.008 0.070
6.0 0.580 0.065 0.036 0.012 0.014 0.070

 

123

substrate towards all the fermentation products. These proportions, shown as percent of
total from C0, are given in Table 22 and Table 23 for the carbon and electron partitioning,
respectively. The fractions of carbon and electrons going from C0 to alcohols
approximately double between a pH of 6.4 and 6.0. Steady-state production of 4-carbon
compounds, i.e., butyrate and butanol, accounts for over 50% of the total electrons from
CO at a fermentation pH of 6.0. Between a pH of 7.2 and 6.0, the partitioning of carbon
and electrons to acetate decreases by approximately 35%. The effects of fermentation pH
on the ratios of 4-carbon products to acetate for this steady-state pH regime are shown in
Figure 26. Even at a pH of 6.0, acetate is the predominant product, on a molar basis, over
all 4-carbon products. Thus, in continuous culture (Grethlein et al., 1990), or in
continuous culture with cell recycling, operation at a pH of 6.0 does not yield butyrate as
the major product compared to acetate, as was observed in batch culture (Worden et al.,
1990).

As stated in the previous section, relative calculations such as the fermentation
stoichiometries and carbon and electron partitioning values of the steady-state fermentation
results are a better means of comparison because of the differences in operating cell
densities, fermenter volumes, and also experimental improvements during the course of
the study. To emphasize these differences, the actual fermentation parameters and steady-
state product concentrations are shown in Table 24. An almost 3-fold variation between
the maximum and minimum steady-state operating cell density was observed, and the
resulting CO consumption rates also varied over this range. Regardless of pH, maximum
product concentrations of 8.7 g/L acetate, 4.6 g/L butyrate, 0.46 g/L ethanol, and 1.2 g/L
butanol were observed during steady-state operation. The corresponding increases in
fermenter productivities, compared to maximum chemostat product concentrations
obtained, are 830% for acetate, 850% for butyrate, 820% for ethanol, 1481% for butanol,
and 2440% for cell mass, assuming 100% viability. Thus, the application of cell recycling

to the continuous fermentation of CO gas by B. methylotrophicwn resulted in a 8-9 fold

124

Table 22. Continuous, steady-state CO fermentation carbon partitioning as a function
of pH for B. methylotrophicum.

 

Carbon Partitioning (%Carbon)
pH 002 CH3COOH C3H7COOH C2H50H C4149OH Cells

 

7.2 55.2 20.5 10.5 1.0 3.2 9.7
6.8 56.2 16.7 12.4 1.2 4.0 9.4
6.4 56.5 15.8 15.2 0.9 3.2 7.0
6.0 58.0 13.0 14.2 2.5 5.5 7.0

 

Table 23. Continuous, steady-state CO fermentation electron partitioning as a function
of pH for B. methylotrophicum.

 

Electron Partitioning (%Electrons)

 

pH CH3COOH C3H7COOH CszOH C4H9OH Cells
7 .2 41.0 26.2 2.8 9.6 20.4
6.8 33.4 31.0 3.8 12.0 19.8
6.4 31.6 38.0 2.7 9.6 14.7

6.0 26.0 35.5 7.5 16.5 14.7

 

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Table 24. Continuous, steady- state CO fermentation parameters, product concentrations,
and CO consumption rates as a function of pH for B. methylotrophicum.

 

Culture Average Product Concentrations (g/L) CO Uptake
pH Volume (L) OD CH3COOH C3H7COOH C2H50H C4H9OH (mol/hr)

 

7.2 1.4 10.3 8.69 3.27 0.315 0.829 0.0218
6.8 1.4 20.6 8.06 4.40 0.459 1.200 0.0360
6.4 1.2 16.3 6.52 4.62 0.282 0.81 1 0.0248
6.0 1.2 7.0 2.71 2.15 0.397 0.698 0.0124

 

Table 25. Continuous, steady-state CO fermentation CO consumption rates as a function
of fermenter volume and culture OD for B. methylotrophicum.

 

Culture CO Uptake CO Uptake CO Uptake CO Uptake
pH Volume (L) (I) (mol/hr) (mol/hr-OD) (mol/hr-L) (mol/hri00D)

 

7 .2 1.4 10.3 0.0218 0.00211 0.01556 0.00151
6.8 1.4 20.6 0.0360 0.00175 0.02573 0.00125
6.4 1.2 16.3 0.0248 0.00152 0.02067 0.00127

6.0 1.2 7.0 0.0124 0.00177 0.01031 0.00147

 

127

increase in acid production and an 8-14 fold increase in alcohol production.

To compare the CO uptake rates of the four steady-state experiments, the rates can be
normalized to both fermenter volume and cell density, as shown in Table 25. The
calculated CO consumption rates for each experiment are shown in both Table 24 and
Table 25. As shown, there is a 2.9 fold difference between the minimum and maximum
CO consumption rates. The effect of fermenter volume can be normalized by calculating
the volumetric CO consumption rates, shown in Table 25. Normalizing the CO uptake
rates by volume, as shown in Table 25, reduces the difference between the minumum and
maximum to a factor of 2.5. Although cell viability was not measured, a speciﬁc CO
consumption rate can be calculated on a basis of overall cell density, and this normalization
results in a 1.4 fold difference between the minimum and maximum uptake rates. A
normalization which takes both volume and cell density into account, with the units for CO
consumption of mol/lmL-OD, is also shown in Table 25, and this calculated uptake rate
results in only a 1.2 fold difference over the range of observed values. These units are
equivalent to the speciﬁc uptake rate (mol/hrgram cell). However, since the viability of
the cells was not measured, the units of mol/hPLoOD will be used henceforth.
Nevertheless, if both fermenter volume and cell density are accounted for, the overall CO
consumption rate during the steady-state fermentations was essentially constant, and thus
the effect of pH on the CO consumption rate between the values of 6.0 and 7.2 can be
assumed to be negligible.

It is important to verify the nature of these steady-state systems with respect to the
assumption of CO mass transfer as the rate-limiting phenomena. The highest value
obtained for the volumetric CO gas consumption was 0.02573 mol/hr-L, or for the
speciﬁc CO gas consumption rate, 0.00439 mol/llrgccn, calculated using a value of 0.344
gceu/l.°OD, calculated from a dry-weight calibration curve (Lynd et al., 1982; Grethlein,
unpub. res.). Calculations from previous batch data (Lynd et al., 1982) have yielded CO
consumption rates as high as 0.02 mol/hr-gccn. The maximum observed speciﬁc CO

128

consumption rate was only 22% of this value, a strong indication that CO mass transfer
was limiting the system during the steady-state cell recycle fermentations.

The mechanism for the predictable metabolic shift towards butyrate production and
away from acetate production as the fermentation pH is lowered has not been determined
experimentally in B. methylotrophicwn. The least realistic hypothesis for the pH effect can
be likened to a deacidiﬁcation mechanism, where the cell responds to an acidic
environment by producing those products which will acidify the environment less. Such
products are alcohols or butyrate, which on a carbon-mole basis has only half of the
acidiﬁcation value of acetate. Yet such a simplistic view of an enormously complex

regulatory process is somewhat misleading, particularly for B. methylotrophicwn, where

the energetics of CO metabolism are intimately linked to the proton-motive force and the

maintenance of a pH gradient across the cell membrane. Many of the current ideas
concerning the effect of pH on metabolic regulation, product formation, and cell growth
arise from studies with heterotrophic, acidogenic bacteria which derive a large proportion

of their metabolic energy from substrate-level phosphorylation. Organisms such as C.

acetobutylicum and C. thermoaceticum are notewothy examples. In the case of B.
methylotrophicum, the production of acetic acid from C0 has two major consequences,
one which centers on the effect of producing acetic acid and ultimately lowering the pH of
the environment, and one which centers on the necessity of acid production for
maintenance of cellular energetics and growth. The former is the topic of discussion in this
chapter, while the latter is discussed primarily in Chapter 5, where the metabolic model
attempts to account for the energetics of the continuous cell recycle fermentations.

The intracellular alkalinity of B. methylotrophicum is required for a pH gradient across
the cyt0plasmic membrane to exist and for various transport and membrane linked
energetic‘functions to operate. Mechanisms such as electron-transport, ATP-linked proton
pumps, and product excretion all participate in the maintenance of the pH gradient, and can
encompass a large proportion of the cells metabolic energy expenditure (Kashket, 1985).

 

 

 

129

The balance of this gradient is paramount to the cell viability, and disruption of the internal
pH frequently results in cell death. For B. methylotrophicum, maintenance of the internal
pH requires that acidic fermentation products such as acetic and butyric acid be excreted
from the cell, along with their requisite protons.

The passive diffusion of most short chain organic acids in chemotrophic, anaerobic

bacteria, has been discussed (Thauer et al., 1977), and the accepted mechanism is that the
un-ionized, or protonated form of the acid is the species which diffuses across the
membrane. The excretion of fermentation products is thus a major mechanism in
controlling intracellular alkalinity. Asuming an initial extracellular environment with a pH
of 7.0, and B. methylotrophicwn with a pH of 6.0 for an intracellular environment, the
production of acetate during growth on C0 would result in diffusion of acetic acid across
the membrane and into the extracellular environment, resulting in a pH decrease over time.
Since the pKa of acetic acid is 4.76 (butyric acid is 4.82), the proportion or ratio of
unionized to ionized species in an aqueous media can be calculated by the equation
ratio = 10(pKa - pH)
where at a pH of 7.0, the ratio is 0.0058, at a pH of 6.0 the ratio is 0.058, at a pH of 5.5
the ratio is 0.182, at a pH of 5.0 the ratio is 0.575, and of course the ratio is 1 at a pH of
4. 76. As calculated, a higher proportion of undissociated species accumulates for acetic
acid as the pH nears 5.0, and this pmportion can freely diffuse back into the cell, thereby
disrupting the pH gradient. 7
The exact pH at which the system allows diffusion back into the cell is a function of the
ref] physiology and the requisite internal pH, for C. thermoaceticum, below a pH of 5 the
remical and electrical potentials across the cell membrane collapse (Baronofsky et al.,
’84). For B. methylotrophicum, the value has not been determined, but based on this
echanism can be assumed to be between 5 and 5.5, from the results observed in the
ntinuous cell recycle fermentations. If the driving force for acetic acid excretion is

mpromised at low pH, and it follows that a logical cellular response to a shift down in

130

the extracellular pH would be to produce less acid, and if acid must be produced, to
produce less acid per mole of substrate.

This somewhat speculative mechanism for the pH effect on the CO fermentation is
supported by the experimental data obtained from batch, continuous, and continuous cell
recycle fermentations. However, it does not explain why at low pH the cell does not cease
acid production altogether, and switch to alcohol production, as is observed during the
solventogenic phase of C. acetobutylicum. This question is addressed by an energetic
analysis of the fermentation. Interestingly, if the concentration of undissociated acid is the
key regulatory factor which inhibits cell functions at low pH, then the pH effect could
perhaps be enhanced by continuous extraction of acids from the fermentation broth,
particularly during low-pH operation. Another hypothesis arising from this analysis is that
the acid uptake results described in Chapter 3 were as much a consequence of the high
initial pH as they were of the high acid concentration, since at equilibrium, 10 g/L at a pH
of 7.2 results in an external acetic acid concentration of approximately 0.04 g/L, while 2
g/L at a pH of 5.0 results in an external acetic acid concentration of 0.73 g/L. If the
external concentration of undissociated acid is the critical species, then the latter case of
this calculation would be more harmful to the cell.

Wallis

The fermentations conducted at pH values of 5.75 and 5.5 initially exhibited primary
butanol production, followed by prolonged oscillations in acetate and butyrate formation.
The time-course proﬁles for these unsteady-state systems are shown in Figure 27 and 28
for the pH values of 5.75 and 5.5, respectively. The oscillations in the pH of 5.5 case are
much more distinct, and oscillations in butanol production are also observed. The initial
responses of primary butanol production are signiﬁcant in that they demonstrate a
capability in B. methylotrophicum for producing butanol as the major product from C0
metabolism. However, these responses were transient and were not maintained for over

86 hours in either the pH of 5.5 or 5.75 case.

.ﬁ

131

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133

Although these initial responses of primary butanol production were only transient
responses, they are evidence for a potential alcohol fermentation from C0 gas. Table 26
shows the initial fermentation stoichiometries for these two unsteady fermentatiOns during
the transient period of primary butanol production. The butanol coefﬁcients in the balanced
equations are signiﬁcantly greater than those in Table 20 for the steady-state experiments.

Table 26. Unsteady-state, initial fermentation stoichiometries as a function of pH for cell
recycle fermentation of CO gas.

 

C] H E 'S'l'

5.75 4co ...> 236002 + .126CH3COOH + .o74c3n7coorr + .021C2H50H + .115C4H90H + .595caus

55 400 -..> zaocoz +.112CH3C001-l + .049C3H7COOH + .029C2H50H + .149C4H90H + ssacaus

 

Averages of the product concentrations over the time period of primary butanol production
were used to obtain these balances. As can be seen in both Figure 27 and Figure 28, the
butyrate concentration was decreasing during the initial phase of each experiment. The rate
of decrease in both cases is approximately equal to the theoretical washout of the acid,
which can be calculated at each data point by assuming complete termination of the
butyrate pathway. The resulting washout curves at each data point, as shown in Figure 29
for the experiment at a pH of 5.5, can be compared to the actual butyrate concentrations;
concentrations above a given washout curve indicate butyrate production, and
concentrations below the washout curve indicate butyrate consumption. As shown, all of
the examined data are on or slightly above each of the washout curves, which is a strong
indication that butanol production occured solely due to CO uptake as opposed to butyrate
uptake.

134

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135

Time zero in Figures 27 and 28 is the point at which the pH had reached setpoint,
approximately 2 days after it had been changed from a value of 6.8, which was the
standard pH used for system start-up and cell growth. Therefore, the transient period of
butanol production from C0 gas, although a real phenomena, could be interpreted as a
result of the experimental procedure, as opposed to a response to fermentation pH. It is
reasonable to assume, with the results obtained in Chapter 3, that the high concentrations
of both acetate and butyrate, resulting from the system start-up at high pH, were an
inﬂuential variable in the observed transient butanol production.

When compared on a basis of carbon and electron partitioning dining this initial time
period, as shown in Table 27 and Table 28, butanol alone accounts for as much as 44% of
the total available electrons in the CO feed at a pH of 5.5. The relative increase in the cell
mass term is not entirely clear. One hypothesis is that cell death is signiﬁcantly greater at
pH values below 6.0. This hypothesis is supported indirectly by the results indicating that
a viable culture could not be maintained at pH values below 5.5, i.e., at 5.35 and 5.1.

The oscillations observed for acetate and butyrate production at low pH suggest that the
energetics of CO metabolism may be a controlling factor during these fermentations, since
formation of these products results in ATP synthesis. Table 14 in Chapter 3 shows the
theoretical net substrate-level ATP and electron yields for each product from C0. These
values were obtained from known and proposed metabolic pathways for CO metabolism
in B. methylotrophicum (Grethlein et al., 1991b). One hypothesis to explain the
oscillatory phenomena is that production of acids and alcohols is regulated by the available
cellular energy (ATP), and only when the ATP levels in the cells are in excess of a critical
concentration does signiﬁcant formation of products other than acetate occur. If such a
hypothesis were true, then maximizing the cellular energetics during CO metabolism,
possibly by using more reduced co—substrates, could alleviate metabolic requirements for
acetate production. The construction and analysis of a metabolic model for CO

fermentation in B. methylotrophicum is the subject of Chapter 5, and the use of co-

136

Table 27. Continuous, initial unsteady-state CO fermentation carbon partitioning as a
function of pH for B. methylotrophicum.

 

Carbon Partitioning (%Carbon)
pH 002 CH3COOH C3H7COOH CszOH C4H9OH Cells

 

5.75 59.0 6.4 7.4 1.0 11.5 14.9
5.5 56.2 5.6 4.9 1.5 14.7 13.3

 

Table 28. Continuous, initial unsteady-state CO fermentation electron partitioning as a
function of pH for B. methylotrophicum.

 

Electron Partitioning (%Electrons)
pH CH3COOH C3H7COOH C21150H C4H9OH Cells

 

5.75 12.6 18.5 2.2 34.5 31.3
5.5 11.2 12.2 4.4 44.1 28.0

 

137

-substrates such as acetate, are described below.

W

The death regime at pH values of 5.35 and 5.1 was indicated by steady washout of all
fermentation products after achieving the pH setpoint in the fermentation system. These
studies were therefore of no practical value, except in that a time-dependent analysis of the
washout curves was the ﬁrst evidence, albeit indirect, for acetate and butyrate uptake
mechanisms in B. methylotrophicum.

W

The hypothesis of energetic limitations during CO fermentation at low pH and the
results for acid uptake presented in Chapter 3 prompted a preliminary investigation as to
the potential for use of cosubstrates such as acetate or methanol concurrent with CO gas.
The purpose of using acetate was to inhibit acetate formation during CO metabolism, and
thus regulate the metabolism more towards butyrate, as was observed in the batch bottle
experiments. This fermentation was conducted with an external acetate concentration of

5.7 g/l at a pH of 5.9, and steady-state was achieved. The fermentation stoichiometry was,
4C0 ---> 2.31C02 + .192CH3COOH + .197C3H7COOH + .037C2H50H

+ .029C4H90H + .330Cells.
which shows that in the presence of acetate, butyrate is the primary product from C0. The
carbon and electron partitioning is shown in Table 29. Although alcohol production was
not affected signiﬁcantly relative to the previous steady-state results, the percent of
electrons contained in 4-carbon products during this fermentation exceeded 57%, which is
the highest value ever observed for this organism growing on C0 gas. This fermentation
was typical in all other respects, with a normalized CO consumption rate of 0.00138
mol/hr-L-OD, which falls approximately at the average of the other steady-state values

calculated. CO conversion for this fermentation was 17.8%.

138

Table 29. Continuous, steady-state CO fermentation carbon and electron partitioning at a
pH of 5.9 in the presence of acetate for B. methylotrophicum.

 

 

(%Carbon) (%Electrons)
Product:
C02: 57.8 0.0
CH3COOH: 9.6 19.2
C3H7COOH 19.7 49.2
C2H50H 1.9 5.6
C4l-190H 2.9 8.7
Cells 8.2 17.4

 

4.4. Summary Discussion of Fermentation Results

The key limitation of the cell recycle studies was the reliance on the carbon and
available electron balances for calculation of CO consumption and cell mass production.
One possible source of error which was calculated not to be a signiﬁcant factor was the
possibility of product stripping by the CO gas leaving the system, a conclusion determined
using equilibrium calculations (Rauolt's Law) based on the typical concentrations of
soluble products and ﬂow rates observed during the fermentations. Besides analytical
techniques, the other most likely source of error was in the measurement of liquid and gas
ﬂow rates, which was done using calibrated pumps and ﬂow meters. These issues must
be kept in mind when discussing the fermentation results.

One interesting set of calculations based on acid and alcohol production from C0 are
the free energies of formation, or 1300' values, for the theoretical reactions of acetate,
butyrate, ethanol, and butanol production. these values are shown in Table 30, where the
equations are balanced in carbon and oxygen only. As shown, all of the formation
reactions are sufﬁciently exergonic to drive ATP synthesis, which has a theoretical energy

requirement of 7.3 kcal/mol.

139

Table 30. Standard free energies of reaction for CO conversiona.

 

 

Reaction AGO' (kcal/g-mol CO)
4CO + 2H20 ---> 2C02 + CH3COOH -9.8

lOCO + 4H20 ---> 6C02 + C3H7COOH -9.7

12CO + 5H20 ---> 8C0; + C4H9OH -9.0
6CO + 3H20 ---> 4COz + C2H50H -8.1

 

aFrom Worden et al., 1991.

Several conclusions relating directly to the feasibility of the coal-derived synthesis gas
conversion concept were drawn during the cell recycle fermentation study. The use of cell
recycle increased overall cell densities 10-20 fold and increased product concentrations 5-
10 fold, compared to continuous fermentations without cell recycling. Since the previous
non-cell recycling fermentations and these studies were all conducted at the same dilution
rate, the corresponding increases in the volumetric productivities of the cell recycle
fermentations was also 5-10 fold. Long-term (6—8 weeks) continuous operation was
achieved without signiﬁcant changes in operating conditions due to membrane fouling or
life. Alcohol production during CO metabolism was measured at all pH values tested, and
a capacity for primary butanol formation directly ﬁom CO gas was observed at pH values
below 6.0. Three metabolic regimes were observed due to fermentation pH, with stable,
steady-state operation at pH values of 6.0 and above, oscillations at pH values of 5.75 and
5.5, and cell death at pH values below 5.5. These ﬁndings have strengthened the

feasibility of an anaerobic fermentation process based on coal-derived synthesis gas.

CHAPTER 5

METABOLIC MODELING AND FERMENTATION ANALYSIS

5.1. Objectives and Principles of the Model

Results from both the metabolic tolerance studies in Chapter 3 and the cell recycle
fermentation studies in Chapter 4 suggested that regulation of CO metabolism in B.
methylotrophicum, speciﬁcally in the production of acetate relative to butyrate and
alcohols, was a function not only of pH but perhaps also of the energetic limitations. To
produce any compounds besides acetate, as shown previously in Table 14, the cell must
expend energy on a substrate-level basis. The relationship between low pH inhibition of
acetate production and loss of ATP generation from decreases in acetate production
suggest that these two antagonistic factors are the key mechanisms which ultimately
determine product stoichiometry during CO fermentation. Thus, in order to more fully
analyze the data obtained during the cell recycle fermentations, the energetics of CO
fermentation in B. methylotrophicum must be taken into account.

From a teleological standpoint, a bacterial cell has the main purpose of performing
work, in the form of transport, motility, and biosynthesis processes, in order to reproduce
in and colonize an environment. The energy needed to support these functions comes from
the metabolism of substrates obtained from the external environment. The efﬁciency of
energy extraction during cell growth on chemical substrates is of paramount importance to
a microorganism's ability to survive, particularly in a competetive environment.

The energy requirements for the physiological functions listed above are met by the use
of energy-rich intermediates, primarily in the form of ATP. Several terms are frequently
used when considiring this energy metabolism, including the cell yield, Ys (g cells/mol
substrate), the ATP yield , YATp (g cells/mol ATP produced), and the ATP gain, GATP

140

141

(mol ATP/mol substrate). However, the calculation of Y ATP requires a knowledge of the
ATP gain, since
YAT'P = (Y s/GATP)

and this raises the problem of experimentally determining these values. Ys can be obtained
by plotting substrate consumption versus the dry weight of cells produced, and taking the
inverse of the slope; for B. methylotrophicum, this value has been determined as 3.4 g
cells/mol CO (Lynd et al., 1982). If G ATP is known, then YATp can be calculated by the
above equation. Typically, a value of 10.5 g/mol is assumed for most anaerobic bacteria, a
value which has been found to be relatively invariant in a wide range of species (Bauchop
and Elsden, 1960). However, this value is based on heterotrophic metabolism with
glucose as the energy source, and cannot be assumed to be accurate for other substrates,
particularly l-carbon compounds (Stouthamer, 1979).

This fact becomes clear when it is realized that values of Y ATP are calculated based on
known metabolic pathways, typically for glucose, and thus the calculation of Y ATP is
dependent on the known production of ATP per mole of substrate. Speciﬁcally, and as
discussed previously, there are two main contributing mechanisms to the generation of
ATP in a bacteria, SLP and ETP. Since the SLP portion of energy generation is that used
to calculate Y ATP. values for Y ATP cannot be readily obtained from organisms with no
net energy generation from SLP, such as B. methylotrophicum. The question then arises
as to how to analyze such organisms from an energetic standpoint, without such key
information. In general, this problem prevents deﬁnite conclusions about the magnitude of
Y ATP in such systems (Stouthamer, 1988). This point has to be considered in the

development of an energetic equation for cell mass production, as described below.

5.2. Development of the Metabolic Model
The use of a fermentation equation to describe the energetics of cell formation during

cell metabolism has been suggested as a means of verifying and analyzing fermentation

142

product data and associated parameters such as maximum theoretical product yields
(Papoutsakis, 1983; Papoutsakis, 1984; Papoutsakis and Meyer, 1985a; Papoutsakis and
Meyer, 1985b; Vallino and Stephanopoulos, 1990). Essentially, these "metabolic
modelling" equations are balanced equations for carbon, ATP, and reducing equivalents,
or electrons, derived from known metabolic pathways (Papoutsakis, 1983). The
application of this type of analyses is that by measuring product formation during
fermentation, the various extents of intracellular metabolic reactions and their selectivities
can be calculated, as a function of only measured product concentrations and ‘cell mass
production.

It follows that in this type of fermentation analysis, one of the most signiﬁcant
reactions is the equation relating the production of cell mass to the substrate, ATP, and
electron requirements. The stoichiometric equation for biomass production requires
knowledge of the cell yields for the particular substrate, a knowledge of Y ATP. and also
calculation of the available electron content of the cell mass and the substrate, so that a
stoichiometric coefﬁcient for the required electrons can be obtained. The ﬁnal result is a
lumped equation for biomass production which relates molar substrate, electron
equivalents, and ATP requirements to cell mass formation.

The method stipulates that the fermentation equation for cell mass is incorporated into a
system of linear equations for the known metabolic pathways. The rates of the product
forming reactions are knowns, the concentrations of intermediate species are assumed to
be constant using a pseudo-steady state assumption, and the system of equations is then
solved for each of the speciﬁc reaction rates. Such a system, depending on the number of
known rates, can be solved uniquely, or, in the case of excess fermentation data, can be
overspeciﬁed, and the best-ﬁt solution to the system of equations obtained using all of the
data. The constraints imposed by the ATP and electron balances used in deriving the
fermentation equations allow maximum theoretical yields and rate-limitin g reactions to be

elucidated, in what is known as identifying "gateway sensors" for metabolic pathways.

143

Application of this type of metabolic modelling to the CO fermentations described in
Chapter 4 is a method by which the energetics and ﬂow of carbon and electrons during CO
metabolism can be studied. However, the method, developed using well-elucidated
heterotrophic bacterial fermentation systems (Papoutsakis and Meyer, 1985a; Papoutsakis
and Meyer, 1985b), has several limitations when applied to the unicarbonotrophic
fermentations by B. methylotrophicum. These limitations arise from the lack of complete
fermentation data, mainly experimental data for cell mass production and CO
consumption, and also from the lack of information on the energetics of CO metabolism,
speciﬁcally to a value for Y ATP- Thus, in order to apply the metabolic model to the
present system, several assumptions and estimates had to be incorporated into the model.

During the following derivation and subsequent analysis, all of the known electron
carriers in the CO metabolic pathway are referred to as NADH, for the purposes of
convenience. This simpliﬁcation is not intended to imply that NADH is the electron carrier
in all of the reactions steps, since some of the speciﬁc electron carriers have yet to be
determined, or are rubredoxin or ferredoxin based (Saeki et al, 189a; Saeki et al, 1989b;
G. J. Shen and J. G. Zeikus, unpublished results). Since the fermentation equations
depend only on the electron balance and not on the speciﬁc carrier, this lack of information
was inconsequential from the standpoint of the model development.

The ﬁrst step in the metabolic modelling of the CO fermentations was to determine an
equation for biomass production from C0. The absence of net substrate-level ATP
production made direct calculation of Ys impossible. As such, the typical form for the
biomass equation,

Substrate + leADHz + yzATP ---> CH1,300,5N0,2
was not applicable, since the molar ATP yield coefﬁcient y2 was unknown. However,
based on the known metabolic pathways, speciﬁcally for acetogenic metabolism, it was

possible to redeﬁne the biomass equation to the form,
Acetyl-CoA + y1NADH2 + yQATP --> 2Cell Mass

144

where the coefﬁcients y1 and y2 were to be determined empirically. In this form, the units
of the yield coefﬁcients are thus mol NADH2/mol Acetyl-CoA consumed for y] and mol
ATP/mol Acetyl-CoA consumed for y2.

Determination of y] and y2 was accomplished using CO growth phase data during
batch bottle experiments (Lynd et al., 1982). This data resulted in a fermentation
stoichiometry of

4CO --> 2.17C02 + 0.74CH3COOH + 0.45Cells.

From the raw data, an electron balance was used to calculate y1, which was 1.48 mol
NADH2/mol Acetyl-CoA consumed, and which was rounded to 1.5 mol NADszmol
Acetyl-CoA consumed. Calculation of y2 required an assumption of ETP and the
stoichiometry of the proton-translocating ATP synthase. The electrons produced by the
oxidation of CO to C02 were assumed to drive an electron transport chain as suggested by
Zeikus et al, (1985), and a standard ratio of 3 moles protons per mole ATP produced was
assumed (Niedhardt, 1987 ). The ATP production via ETP could thus be determined using
Lynd's data. From the known SLP contribution, taken from the pathway of CO
metabolism and acetate formation (See Figure 5), the net overall ATP consumption, and
thus the ATP requirement for cell mass production, was then calculated. This value was
found to be 2.0 mol ATP/mol Acetyl-CoA consumed, which interestingly results in a Ys
value of 0.5 mol ATP/mol CO, or a YATp value of 6.8 g cell/mol ATP.

A system of linear equations was set-up based on the developed pathway for CO
metabolism in B. methylotrophicum, but including pathways for ethanol and butanol
formation (J. S. Shieh, unpublished results). A detailed diagram of this pathway,
including the hypothesized ﬂow of electrons during electron transport phosphorylation, is
shown in Figure 30. This scheme is an extension of that suggested by Zeikus et al,
(1985). It must be emphasized that the actual electron carriers in some of the reactions

have not yet been deterrriined, although NADH is used as a generic carrier in Figure 30.

145

Illllm‘t'IIllIJr Membrane Extracellular

2C() 31”)

 

—-‘ 4C0

ADP.P..HO + [ED
' 2 2NAD N

+ HC‘JTHF 20,,
211
H o ZNADH
2 + +
tic-=- F - > 2“

 

 

 

 

 

 

 

 

 

 
 
 

 

- 256
HZC- 2m ‘J__N
[C01
#
2F
H3C—THF
ADP + I: +
&__.ll:ml= <— - "' — "‘3“
HS-CoA I C FR AT
cuscocoa a- - 1 ‘3 , CH30H
HS-C A ‘ ‘ I C I
° ' \ ' R” H o
t‘ H CoA 2
c0 H coou————>
o (0 CH 3c H3CO- c 3 CH3COOH
CH3'éc Ecol». ADP ATP
NADH+ + +
+ NADH + H HS‘C0A
NAD
- 9 + 9
C —
H3CH(OH)CF5C-C0A 3CH CH3CH20H CH3CH20H
+ +
1120 O NADH + H NAD
II

CH 3CH CHC-CoA

NADH“!+ c c c g c c coon COOH
; P. "3155' 7 S'G's'hﬂz 050101
"AD 9 HS-COA ADP ATP

l
CHBCHZCHZC-Co

 

+
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0 .
X” 113CH2Cl-12' ' enacnzcuzcuzolr crgcrgcrgculon

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Figure 30. Known and proposed metabolic pathways for carbon and electron
flow during C0 metabolism in B; methylotrophicum.

 

146

Letters indicate unknown electron carriers, with F indicating ferredoxin and R indicating

rubredoxin. The equations were developed from each of the major reaction steps, where
CO --> C02 + NADHz (l)

is the reaction for oxidation of CO to C02. For the 6-step tetrahydrofolate—mediated

conversion of the bound C02 group to the bound CH30H group, or CH30-, an overall

reaction is

C02 + 3NADH2 + ATP --> CH30- (2),
and for the condensation reaction to form acetyl-CoA,

C0 + CH30° --> Acetyl-CoA (3),
is the carbon-balanced equation.

Acetyl-CoA participates in several pathways, including the formation of acetate,with
the overall reaction

Acetyl-CoA ---> CH3COOH + ATP (4),
and for the formation of ethanol,

Acetyl-CoA + 2NADH2 ---> C2H50H (5),
and for the formation of cell mass,

Acetyl-CoA + 1.5NADH2 + 2ATP --> 2Cell Mass (6),
and the formation of acetoacetyl—CoA,

Acetyl-CoA -->0.5Acetoacetyl-CoA (7).
Acetoacetyl-CoA is the precusor to butyryl-CoA in a 3-step reaction sequence simpliﬁed
by the overall reaction

Acetoacetyl-CoA + 2NADH2 --> Butyryl-CoA (8)
which then can form either butyrate by

Butyryl-CoA ---> C3H7COOH + ATP (9)

or butanol by
Butyryl-CoA + 2NADH2 ---> C4HgOH (10)

147

The complete system of fermentation equations for this pathway, balanced in carbon,

ATP, and electrons (NADH), is thus
CO ---> C02 + NADH2

(1)

C02 + 3NADH2 + ATP --> CH3O~ (2)

C0 + C1130 ---> Acetyl-CoA

(3)

Acetyl-CoA ---> CH3COOH + ATP (4)
Acetyl-CoA + 2NADH2 --> C21-IsOH (5)
Acetyl-CoA + 1.5NADH2 + 2ATP ---> 2Cell Mass (6)

Acetyl—CoA --->0.5Acetoacetyl-CoA ' (7)
Acetoacetyl—CoA + 2NADH2 ---> Butyryl-CoA (8)
Butyryl-CoA ---> C3H7COOH + ATP (9)
Butyryl-CoA + 2NADH2 ---> C4H90H (10)

If the rates of the ten reactions l-10 are given as a-j, then all of the reactants, products,

and intermediate concentrations are thus a function of these 10 equations, where

N ADHz =
CH30- =
Acetyl-CoA =
CH3COOH
C2H50H =

Cell Mass =
Acetoacetyl-CoA =
Butyryl-CoA
C3H7COOH =
C4H90H =

(a-3b-2e-2f-2h-2j)
(b-C)

(c-d-e-f-g)

(d)

(e)

(20

(0.5g-h)

(h-i-J)

G)

(i)

and where the ATP balance, because of the presence of ETP, was not used to specify the

system.

148

Several pseudo-steady state assumptions were then invoked on all the intracellular

reaction intermediates, leaving 5 independent equations:

NADH2 : (a-3b-2e-l.5f-2h-2j)=0
C1130 : (b - c) = 0
Acetyl-CoA : (c-d-e-f-g)=0
Acetoacetyl—CoA : (0.5g - h) = 0
Butyryl-CoA : (h - i - j) = 0.

The measurable or calculated species, acetate, butyrate, butanol, ethanol, C0, C02 , and
cell mass, comprise the other speciﬁcations, resulting in a system of 10 total equations
with 5-independent and 5-dependent equations. If all of the species concentrations were
measured, this system could conceivably be overspeciﬁed with 12 total knowns or
equations (7 measurable rates and 5 independent equations) and 10 unknowns (rates a-j).
Unfortunately, depending on the CO fermentation system, only 5 measurable quantities
were available in any given experiment, and thus the system is uniquely speciﬁed. Yet
some of the unmeasurable species can be calculated from the carbon and available electron
balances, to overspecify the system.

This issue is addressed for each of the speciﬁc cases of the model, applied below to the
chemostat experiments conducted previously in order to verify the overall approach. The
model is then applied to the steady-state cell recycle experiments using different
speciﬁcations, and then also in a prelimary approach to modelling the unsteady-state cell

recycle experiments.

5.3. Application to Chemostat Results

The model developed above was ﬁrst applied to the previously obtained steady-state
chemostat data in order to verify the approach and to determine some of the energetic
parameters which are possible to calculate using the metabolic model. The system of

equations described above was solved with data for acetate, butyrate, ethanol, butanol,

149

and cell mass production, and with the 5 independent equations; it was thus a unique
system of equations with 10 knowns and 10 unknowns. The method of solution was
Gaussian elimination, or row-reduction techniques. The following solutions were
obtained, where rAc, rBu, rEt, rBh, and rCM are the measured rates of acetate, butyrate,
ethanol, butanol, and cell mass production, respectively:

a = 3[rAc + rEt + 0.5rCM + 2(rBu + rBt)] + 2rEt + 0.75rCM + 2(rBu + rBt) + 2rBh

b = rAc + rEt + 0.5rCM + 2(rBu + rBt)

c = rAc + rEt + 0.5rCM + 2(rBu -l- rBt)

d = rAc e = rEt f = 0.5rCM g = 2(rBu + rBt) h = rBu + rBt

i = rBu j = rBh.

A spreadsheet was then used to apply the chemostat fermentation data to these
solutions. All of the rate data, as a function of pH, were normalized to the cell density for
these calculations, which thus yielded reaction extents. Some of the important calculations
were the molar consumption and production extents for CO and C02, shown in Figure
31. These extents were relatively constant over the tested range of pH values. As a check
on the validity and accuracy of the model, the ratio of CO consumption to C02 production
calculated using the model can be compared to the molar ratios calculated using carbon and
available electron balances, as were shown previously in Table 17. This comparison is
made in Table 31. As shown, there is less than a 6% difference in all of the calulated
values. The model was thus veriﬁed and the accuracy, relative to the carbon and electron
balancing method, was high. Other important calculations were the ratio of acetyl-CoA
derived products to butyryl-CoA derived products, which includes cell mass, and also the
ratio of 2-carbon products to 4-carbon products, which does not include cell mass. These
values are shown in Figure 32, where as seen previously, almost a 7-fold decrease in the
2-carbon to 4wcarbon product ratios occurs between a pH of 6.8 and 6.0. Finally, the
energetics of CO metabolism were calculated by assuming ETP based on the pathways

shown in Figure 30. A ratio of 3 moles of protons consumed to every mole of ATP

CO Consumption (mol CO/mol Cell Mass)
(D2 Production (mol C02/mol Cell Mass)

Molar Product Ratio

150

 

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5.40 5.60 5.80 6.00 6.20 6.40 6.60 6.80 7.00
Fermentatioan

"I'I'Il'

 

 

 

Figure 31. Molar consumption and production rates for CO and
CO2 gas as a function of fermentation pH during

steady-state, continuous, culture.

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Fermentation pH

 

 

 

Figlne 32. Molar product ratios as a flmction of fermentation pH
during steady-state, continuous culture.

151

Table 31. Comparison of CO:COz ratios from chemostat culture between metabolic
model and carbon and electron balancing calculations.

 

 

CO:C02 ratio
Metabolic Model Carbon and Electron Percent
Culture pH Calculation Calculation Difference
6.8 1.805 1.914 5.69
6.5 1.792 1.878 4.58
6.0 1.704 1.762 3.29
5.5 1.749 1.835 4.69

 

produced was assumed (Niedhardt, 1987), and the protons calculated by the reaction
extent of (b),
C02 + 3NADH2 + ATP --> CH30°.

The relative theoretical ATP contributions calculated for ETP, SLP, and also the total
theoretical ATP production for the chemostat fermentations, normalized to the total CO
consumption, are shown in Figure 33 as a function of pH. As has been described
previously, substrate-level reactions consume ATP during production of all products other
than acetate, as observed by the negative contribution of SLP in Figure 33. However, with
the calculated contribution of ETP, the net ATP generation during CO metabolism was
positive, and did vary somewhat with pH. This method of energetic analysis is therefore
consistent with the hypothesis that during autotrophic or unicarbonotrophic growth, ATP
deﬁcits during catabolism are compensated for by an active ETP mechanism, which results
in positive ATP production and thus cell growth.

Overall, the results of the test of the model on the chemostat data were encouraging,

with all of the calculated values for CO consumption, C02 production, product ratios, and

energetics being consistent with either previous calculations or current theory. It was thus

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determined that this method of metabolic modelling was applicable to the analysis of the

cell recycle fermentations, as described below.

5.4. Application to Steady-State Cell Recycle Fermentation Results

The application of the metabolic model to the steady-state continuous cell recycle
fermentation results was different from the application to the chemostat data in several
ways. First was that instead of cell mass, C02 production was the fifth measured variable.
Second, the variability of the raw data from the cell recycle experiments, as shown in
Table 24, required normalization of the data prior to solution of the system of equations.
Solutions to the rates of reactions are thus best compared in the units of mol/hr-L00D, as
described previously. The steady-state results were analyzed in 5 separate cases. Case 1
was a unique solution to the system of equations but without the addition of the
fermentation equation for cell mass production, i.e., the reaction rates for only
extracellular products were obtained. Case 2 is the so-called base-case, and should be
considered as the representative case for this system. In Case 2, the system was uniquely
speciﬁed, and CO consumption and cell production calculated. Case 3-5 were applications
of calculated data from the carbon and available electron balances in order to overspecify
the system. In Case 3, the system was overspeciﬁed with the calculated value for C0
consumption. In Case 4, the system was overspecified with the calculated value for cell
mass production. In Case 5, both the value for CO consumption and the value for cell
mass production were used to overspecify the system.

SHE 1.11. 15 'ﬁlS 11:11” 131°

In Case 1, the system was set-up without an equation for the production of cell mass
from acetyl-CoA. This case was evaluated to determine the contribution, or lack of
contribution, of cell production on the overall energetics of the CO fermentations. The

system of equations used was

co ---> 002 + NADHZ (1)

154

C02 + 3NADH2 + ATP ---> CH30° (2)
00 + CH30- ---> Aeetyl-CoA (3)
Acetyl-CoA ---> CH3COOH + ATP (4)
Acetyl-CoA + 2NADH2 ---> C21150H (5)
Acetyl-CoA -->O.5Acetoacetyl-CoA (6)
Acetoacetyl-CoA + 2NADH2 --> Butyryl-CoA (7 )
Butyryl-CoA ---> C3H7COOH 4» ATP (8)
Butyryl-CoA + 2NADH2 --> C41190H (9).

and since there are only 9 equations with solutions a-i, the independent relationship
NADHz = (a - 3b - 2e - 2g - 2i)

was not used to solve the system. Instead, the system was solved with the remaining 4
independent equations, and the 5 measured variables rAc, rBu, rEt, rBh, and rC02.
Gaussian elimination was used to solve this unique system of equations, with solutions of

a = rC02.+ rAc + rEt + 2(rBu + rBt)

b = rAc + rEt + 2(rBu + rBt)

c = rAc + rEt + 2(rBu + rBt)

d = rAc e = rEt f = 2(rBu + rBt) g = rBu + rBt

h = rBu i = rBh.
with calculation of unknowns such as the rate of CO consumption (rCO), the rate of
electron production/consumption (rNADH), and the rate of ATP production via SLP,
ETP, and their combined total (rSLP, rETP, rA'I'P) by the equations

rCO = (a + c)

rNADH = (a-3b-2e-2g-2i)
rSLP = (d + h - b)

rETP = 2/3(b)

rATP = (d + h - b + 2/3(b)).

155

Figure 34 shows the results if the data are not normalized to fermenter volume or cell
density, where large variations in the calculated rates for CO consumption and cell density
are observed. However, if the data are normalized, the CO consumption rates are
considerably smoothed over the pH range, as shown in Figure 35. Subsequently, all of
the data presented will be normalized in this manner. Figure 36 shows the molar
production and consumption of CO and C02, respectively. These ratios are also relatively
constant over the pH range. Also, in Figure 37 are the acetate to butyrate ratio and the 2-
carbon to 4-carbon product ratio as a function of pH. These ratios are both approximately
halved as the pH is decreased from 7.2 to 6.0. In Figure 38 are the molar yields for ATP
generation, where once again, the SLP contribution is negative, but with positive
production overall as a result of the ETP contribution. Importantly, there is also an
observed trend of decreasing ATP production as the pH decreases, which is due to an
increase in the ATP deﬁcit during SLP, as shown by the larger negative values for SLP as
the pH drops from 7.2 to 6.0. This observation, even though Case 1 is not representative
of the actual biological system since cell mass was not accounted for, is an important
aspect of the pH effect, and is discussed in detail for the more appropriate cases below.
Finally, as a check on the accuracy of the metabolic model to the cell recycle data, rNADH
can be broken down into the electron-producing and electron-consuming reactions, as
shown in Figure 39. Since the net production of electrons, or NADH, should be the
electron requirement for cell production, this value should be positive, as is indicated.

The results for Case 1 indicate that the metabolic model, as for the chemostat data, is
consistent when applied to the cell recycle data. Furthermore, as discussed in Chapter 4,
Case 1 has reemphasized the need for normalization of all fermentation rate data to both
fermenter volume and culture density, in order to fully compare the calculated reaction
rates. The production of ATP as a function of fermentation pH was also observed to be

dynamic, and is detailed in a later section.

 

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CO Consumption (mol/hr)
Cell Production (mol/hr)

 

0.010 ‘ +Cdl Production (x10)
0.005pup“.1...n-.r-v1-v-.6.o

6.0 6.2 6.4 6.6 6.8 7.0 7.2
Fermentation pH

 

 

Figure 34. Steady-state CO consumption, cell production, and
optical density as a function of pH for continuous
cell recycle fermentation.

 

 

 

 

 

 

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Figure 35. Normalized CO consmnption rates for continuous
cell recycle fermentation as a function of pH.

(am 099) htsuoa 1939(10

CO Consumption (mol/hr-L-OD)
(302 Production (mol/hr-L-OD)

157

 

 

 

 

 

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Figure 36. Normalized molar gas consumption and production
rates as a function of pH during steady-state,
continuous cell recycle fermentation: Case 1.

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Figure 37. Molar product ratios as a function of pH during

steady-state, continuous cell recycle fermentation:
Case 1.

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For an accurate analysis of the cell recycle fermentation results, the production of cell
mass, or more speciﬁcally the proportioning of carbon and electrons to cell mass, must be
accounted for in the metabolic model. Case 2 includes the derived equation for cell mass

production from acetyl-CoA into the system of equations,

CO ---> C02 + NADH2 (1)
C02 + 3NADH2 + ATP --> CH30- (2)
00 + CH30° --> Acetyl-CoA (3)
Acetyl-CoA ---> CH3COOH 4» ATP (4)
Acetyl-CoA + 2NADH2 --> C2H50H (5)

Acetyl-CoA + 1.5NADH2 + 2ATP --> 2Cell Mass (6)

Acetyl-CoA -->O.5Acetoacetyl-CoA (7 )
Acetoacetyl-CoA + 2NADH2 --> Butyryl-CoA (8)
Butyryl-CoA ---> C3H7COOH + ATP (9)
Butyryl-CoA + 2NADH2 --> C4H90H (10).
with rates a-j for reactions l-10. The 5 independent equations
NADH2 : (a-3b-2e-l.5f-2h-2j)=0
CH30 : (b - c) = 0
Acetyl-CoA : (c-d-e-f-g)=0
Acetoacetyl-CoA : (0.5g - h) = 0
Butyryl-CoA : (h - i - j) = O

are used to solve the system with the rate data rAc, rBu, rEt, rBh, and rCOz. The solution

set for this unique system (10 equations or knowns and 10 unknowns) is
a = rC02.+ 3/7[rAc + rEt + 2(rBu + rBt) + 52]
b = 3ﬂ[rAc + rEt + 2(rBu + rBt) + 0] c = 3/7[rAc + rEt + 2(rBu + rBt) + Q]
d = rAc e = rEt f = 3/79 - 4/7(rAC + rEt + 2rBu + 4rBh)
g = 2(rBu + rBh) h = rBu + rBh i = rBu j = rBh.

161

and where Q = 2/3(r002 - 2rEt - 2rBu - 4rBh), which were all obtained again by
Gaussian elimination.

The molar consumption and production rates for CO and 002 are shown in Figure 40
for the Case 2 results, and with the addition of the cell mass equation there is an increase
in those rates compared to the Case 1 results. To again compare the results of the model
with previously calculated values for CO consumption and 002 production, as shown in
Figure 36, the carbon and available electron balance values, shown previously in Table
20, were used. This comparison is given in Table 32. The accuracy, upon comparison, is

Table 32. Comparison of 00:002 ratios from steady- state cell recycle culture between
metabolic model and carbon and electron balancing calculations for Case 2.

 

 

002002 ratio
Metabolic Model Carbon and Electron Percent
Culture pH Calculation Calculation Difference
7.2 1.747 1.810 3.48
6.8 1.715 1.778 3.54
6.4 1.718 1.770 2.94
6.0 1.683 1.724 2.38

 

excellent, with less than 4% difference in all cases.

The ratio of acetyl-CoA to butyryl-CoA derived products and 2—carbon to 4-carbon
derived products is shown in Figure 41, and these ratios are consistent in that compared to
the results in Figure 37, the ratio of acetyl-CoA to butyryl-CoA derived products has
increased relative to Case 1, while the ratio of 2-carbon to 4-carbon products remains
essentially unchanged. In Figure 42, the molar ATP production is presented, and an
interesting observation is that the total ATP production is positive only because of the ETP
component. Unlike what was observed in Case 1, the total molar ATP production in Case

162

 

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Figure 40. Normalized molar gas consumption and production

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Figure 41. Molar product ratios as a function of pH during

steady-state cell recycle fermentation: Case 2.

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2 is highest at a pH of 6.0 and lowest at the intermediate values of 6.8 and 6.4. The actual
data are, in moles ATP produced/moles CO consumed, 0.0175, 0.0111, 0.0116, and
0.0221 for the pH values of 7.2, 6.8, 6.4, and 6.0, respectively. The molar cell
production of these same experiments was calculated as 0.000076, 0.000046, 0.000048,
and 0.000019, for the pH values of 7.2, 6.8, 6.4, and 6.0, respectively, all in units of
mol cells/hr'LoOD. The fermentation equation for cell mass production, which reflects an
ATP consuming process, can thus be concluded to be an important equation in the cellular
energetics, with decreased cell mass production resulting in a increase in ATP yield. This
result is opposite to the effect of the relative decrease in acetate production as the
fermentation pH is decreased during 00 metabolism in B. methylotrophicum, which
results in a decrease in the ATP yield. Under the conditions of this analysis, the
relationship between these two anatagonistic trends is complicated. Nevertheless, the
coefficient for ATP in the cell mass equation
Acetyl-CoA + 1.5NADH2 + 2ATP --> ZCell Mass (6)

is therefore a critical value which requires careful derivation for accurate energetic
analysis.

The majority of ATP obtained from the ETP mechanisms must be used to support the
catabolic reactions which consume ATP via SLP. This point supports the argument for
cosubstrate addition such as CH30H, which could alleviate the negative effect of SLP on
the overall energetics of CO metabolism, but only in a relative manner, since decreases in
the ﬂow of electrons in the ETP cascade would result from relative decreases in CO
consumption. This effect could well be analyzed by incorporating a cosubsu‘ate reaction
into the model system of equations. The calculated value for total NADH shown in Frgm'e
43, which is multiplied by a factor of 10 in the graph, can be taken as the relative error of
the solution set for this system, since the NADH balance must equal zero by the pseudo-
steady state assumption. This closure of the system was achieved with high accuracy, as

the relative error shown is approximately 3% in all cases.

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An important parametric study which was conducted with the Case 2 system of
equations was the effect of the proton-translocating ATP synthase stoichiometry on the
yields of ATP via ETP as a function of pH, and also on the total yield of ATP. This
stoichiometry was assumed in all cases to be 3 protons for every mole of ATP produced.
The effect of this stoichiometry was investigated by varying the ratio from 2.5-3.0, which
are typical values (Niedhardt, 1987). The results of this study are shown in Figure 44. As
shown, a constant increase in the proton/ATP ratio results in a linear decrease in both the
ETP yield of ATP and also the overall ATP yield, as expected. Since the actual existence
of ETP in B. methylotrophicum has not yet been conclusively proven, particularly as to
the existence of a proton translocating ATP synthase, the choice of the proton/ATP ratio
was not a signiﬁcant factor in the development of this metabolic model.

The use of overspecified systems when solving the fermentation system of equations
can theoretically show the accuracy of the data, by selective overspeciﬁcation comparison
methods. In this study, there were only enough measured variables to completely specify
the system. By using calculated values for cell mass production and CO consumption
shown in Chapter 4, however, it is possible to overspecify the system of equations, and
this is the focus of Cases 3-5. Solution of an overspeciﬁed system cannot be accomplished
by row-reduction methods, but requires a least-squares solution of the form,

[A] [X] = [r]
[x] = (ATA)-1AT[r1
where [r] is the vector set of known reaction rates, [A] a matrix of the system of
equations, and [x] the solution vector a-j.

The same analyses for Case 2 were conducted with Cases 3-5 except that in Case 3, the
calculated value for CO consumption was used, in Case 4, the calculated value for cell
mass production was used, and in Case 5, both calculated values were used. The validity

of using values based on carbon and electron balances of the fermentation data is suspect,

167

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168

since the system of fermentation equations is itself a system of carbon and electron
balances. For example, in Case 3, the speciﬁed ethanol production rate at a pH of 6.4 was
0.000006 mol/hr‘LOOD, but the solution obtained gave a value of 0.000026 mol/hrL-OD.
In Case 4, the speciﬁed ethanol production rate at a pH of 7.2 was 0.000009
mol/hr-L-OD, but the solution obtained gave a value of 0.000085 moUhroL-OD. Other
similar examples indicate that the electron balance in the system of equations was severely
affected.

The overall results from the overspeciﬁed cases are shown in Figures 45, 46, 47, and
48. In Figure 45, the molar C0 consumption and C02 production are shown to be
invariant with respect to Case 2. In Figure 46, the 2-carbon to 4-carbon product ratios for
Cases 3-5 are quite variant compared to Case 2, relatively invariant compared to each other
at low pH, and variant compared to each other at high pH. Figure 47 shows the same
trend, where compared to Case 2, the molar ATP yield is much higher in the overspeciﬁed
cases, a result due primarily to calculated rates of cell mass formation which were much
lower than in Case 2 at pH values above 6.0. As stated previously, the molar electron
yield should equal zero because of the pseudo-steady state assumption invoked in the
development of the metabolic model. Case 2 was shown to be relatively accurate in this
regard. But, as seen in Figure 48, Cases 3-5 are more accurate, particularly at low pH, in
achieving closure of the electron balance, a result due to the anomolous variations
observed in the spread-sheet data. Therefore, although the accuracy of the the results for
Cases 3-5 are good, the solutions for several of the speciﬁed parameters, as discussed
above, where not consistent with the speciﬁcations used to solve the system of equations,

and are thus not representative of the actual solution of that system.

5.5. Application to Unsteady-State Cell Recycle Fermentation Results
The system of fermentation equations used to solve for the steady-state reaction rates is

equally applicable to the unsteady-state results in theory (Papoutsakis, 1984; Papoutsakis

169

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5.8 6.0 6.2 6.4 6.6 6.8 7.0 7.2 7.4
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Figure 45. Molar CO consumption for Cases 2—5 as a function of
pH during steady-state cell recycle fermentation.

   
  
  

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Fennentatioan

Figure 46. Molar product ratios for Cases 25 as a function of
pH during steady-state cell recycle fermentation.

Theoretical Molar ATP Yield (mol ATP/Incl CO)

Molar Electron Yield (mol e’/mol CO)

25>
8
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170

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Figure 47 . Theoretical molar ATP yields for Cases 2-5 as a
function of pH during steady-state cell recycle

fermentation.

 

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Fermentation pH

Figm'e 48. Theoretical molar electron yields for Cases 2-5 as a
function of pH during steady-state cell recycle

fermentation.

171

and Meyer, 1985a), but to do so requires more rate information from the fermentations. A
generic mass balance on a continuous, constant volume reactor as a function of time is
VR-(dC/dt ) = FiCi - FoCo + VRRG
where the subcripts i and 0 correspond to the flows in and out of the reactor, respectively,
and where VR is the reactor volume, C the concentration of the species, F the ﬂow rates,
and R0 the species generation or production term. Deﬁning D, the liquid dilution rate, as
F/V R and assuming the concentration of species C in the incoming stream is 0, then
(dC/dt ) = - DCO + RC,
and at steady-state,
R0 = DCo
which was the equation used to calculate the known rates of acetate, butyrate, ethanol, and
butanol formation for the steady-state analysis above. In the unsteady-state case,
RC, = (dC/dt) - DCO
and to solve this equation requires a value for dC/dt.

In an unsteady-state system, the terms in the above equation can potentially assume
different values at each time point, or in the context of the cell recycle fermentation, at each
data point. Therefore, to apply the system of fermentation equations and solve for the rates
of reaction, not only will the values for dC/dt have to be obtained, but the system of
equations will have to be solved for each separate data point. These complications
distinguish the unsteady-state case from the steady-state case.

To obtain values for dC/dt, the fermentation data for each of the species from the cell
recycle fermentation data at a pH of 5.5 were plotted separately as a function of time,
similar to what was shown in Figure 28. The method of graphical differentiation was then
applied to obtain dC/dt at each data point for each of the 5 species, acetate, butyrate,
ethanol, butanol, and C02. These data, along with the calculated values for DCO at each

point were used to solve the system of equations derived from
C0 --> C02 + NADH2 (1)

172

C02 + 3NADH2 + ATP ---> (II-I300 (2)
C0 + CH30° ---> Acetyl-CoA (3)
Acetyl-CoA ---> CH3COOH + ATP (4)
Acetyl-CoA + 2NADH2 --> C2H50H (5)

Acetyl—CoA + 1.5NADH2 + 2ATP --> 2Cell Mass (6)

Acetyl-CoA -->0.5Acetoacetyl-CoA (‘7)

Acetoacetyl-CoA + 2NADH2 --> Butyryl-CoA (8)

Butyryl-CoA --> C3H7COOH + ATP (9)
Butyryl-CoA + 2NADH2 --> C4H90H (10).

with solutions a- j for equations 1-10, and which was uniquely speciﬁed with the 5

independent equations
NADHz : (a-3b-2e-1.5f-2h-2j)=0
CH30- : (b — c) = 0
Acetyl-CoA : (c-d-e-f-g)=0
Acetoacetyl-CoA : (0.5g - h) = 0
Butyryl-CoA : (h - i - j) = 0

and the rate data rAc, rBu, rEt, rBh, and rCOz at each point. The solution set for this
unique system (10 equations or knowns and 10 unknowns) is the same, since the system

of equations is the same, as for Case 2, described above, namely

a = rC02.+ 3/7[rAc + rEt + 2(rBu + rBt) + Q]

b = 3/7[rAc + rEt + 2(rBu + rBt) + (2] c = 3/7[rAc + rEt + 2(rBu + rBt) + Q]

d = rAc e = rEt f = 3/79 - 4/7(rAC + rEt + 2rBu + 4rBh)

g = 2(rBu + rBh) h = rBu + rBh i = rBu j = rBh.
where Q = 2/3(rC02 - 2rEt - 2rBu - 4rBh), and which were all obtained again by
Gaussian elimination. This solution set allowed calculation, as a function of time, of all the

same fermentation variables.

173

Shown in Figure 49 are the normalized CO consumption and C02 production values.
Immediately it is evident that the fermentation at a pH of 5.5 cannot possibly be mass
transfer limited, since the rate of CO consumption exhibits strong and distinct oscillations
over the entire fermentation time-course. However, the maximum rates of CO
consumption are comparable to those observed in the steady-state cases (Figure 40), on
the order of 0.00125 mol/hr'LoOD. The production of C02 exhibits the same oscillatory
phenomena. The molar yields of both butyrate and acetate were also calculated. Figures 50
and 51 show the molar acetate yield and molar butyrate yield as a function of time,
respectively. Oscillations are also observed in these Values, and thus normalization to the
oscillating CO consumption rate does not dampen the oscillations in acetate and butyrate
production rates. Finally, the theoretical molar ATP yields, normalized to the CO
consumption rates, are shown in Figure 53. The ATP yield from the ETP contribution is
constant over time, which is expected since this contribution is calculated from the single
reaction in the system of equations,

C02 + 3NADH2 + ATP ---> (II-I300 (b)

and since the rate of C02 production is oscillating as a result of the oscillating rate of CO
consumption, normalization of the ETP rate of ATP generation would be expected to
dampen those oscillations. However, as seen in the ﬁgure, the ATP yields from SLP do
show oscillations, and those oscillations are directly reﬂected in the oscillations in the total
ATP yields. This result could be expected from the results seen in Figures 50 and 51,
where the oscillations in acetate and butyrate production are independent of CO
consumption, and thus the normalized SLP contribution, which is largely due to the
contributions from the acid forming reaction steps, would not be dampened by oscillations
in CO consumption.

From Figure 52, the overall theoretical ATP yield is negative over the entire time-
course, which is not consistent with known biological phenomena. However, the

magnitude of the ATP yields is consistent with those shown in Figure 42, particularly for

174

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recycle fermentation at a pH of 5.5.

176

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177

the ETP contribution. Direct comparison with the results in Figure 42 at a pH of 6.0 show
that the average SLP contribution is much less (more negative) at a pH of 5.5 than at a pH
of 6.0, which is consistent with the reduced acetate production observed. It is therefore
reasonable to assume that the negative ATP yield at a pH of 5.5 is not a result of the
model, since the model accurately calculated the magnitudes of the ETP contribution and
CO consumption rates. It is also apparent that the negative ATP yield cannot be a function
of the chosen proton-translocating ATP synthase stoichiometry, since although that was
shown to inﬂuence the contribution of ETP in Figure 44, the inﬂuence of that variable is
not of the same magnitude as the changes observed in Figure 52. The possibility of an
inaccuracy in the fermentation equation for cell mass production,
Acetyl-CoA + 1.5NADH2 + 2ATP --> 2Ce11 Mass (f)

is one potential source of error, although the Case 2 system was solved with a high degree
of accuracy using this equation.

One possible explanation for the negative theoretical ATP yields is that not all of the
electron-transport mechanisms were included in the model. There has been some evidence
(G. J. Shen and J. G. Zeikus, unpublished results) indicating that the electron carrier (X)
for the reaction

Crotonyl-CoA + XH --> Butyryl-CoA
is membrane bound, which is indirect support for its participation in an electron transport
pathway. If true, then the energetics of butyrate production would be much more favorable
than indicated in this study, due to the additional ATP generation from the ETP-linked
step. Other mechanism of ATP synthesis, speciﬁcally solute-transport or monovalent ion-

linked systems are also possible but purely speculative at this time.

5.6. Limitations of the Model
The metabolic model developed for the cell recycle fermentation results has wide

potential in conducting analysis of various fermentation pathways, both from C0 and from

178

other substrates, in B. methylotrophicum. With the long-term nature of the typical CO
fermentation, such a technique is invaluable. However, there are several limitations of this
model, particularly in the assumptions and speciﬁcations used to develop the system of
equations.

It is important to realize that although the energetic analysis was based on known
metabolic pathways, such an analysis is no substitute for actual measurement of species
such as ATP. This limitation is reﬂected in the term "theoretical", which is used as a
qualiﬁcation on the ATP yields presented in this chapter. In general, one can safely
conclude the ATP yields calculated by the model are indicative of the general trends of
ATP charge in the cell, but not reﬂective of the actual species concentrations.

The effect of the cell mass equation on the overall theoretical yields for ATP and
electrons was found to be signiﬁcant, particularly for the Case 2 results, where the ATP
yield was greater at a pH of 6.0, presumably from a decrease in the production of cell
mass. This equation,

Acetyl-CoA + 1.5NADH2 + 2ATP --> 2Cell Mass (6)
requires more detailed development and analysis, both as to its applicability and to the
speciﬁc values for the coefﬁcients for NADHz and ATP. Results which were not shown
have indicated that the coefﬁcients are extremely important for calculation of cell mass
production, since in all cases analyzed, equation 6 was used to solve the electron balance
for NADH2, where
NADH2: (a - 3b - 2e -1.5f - 2h - 2j) = 0,
or speciﬁcally,
(a-3b-2e-2h-2j)= 1.5f

which is why the cell production rates where not highlited in this analysis.

Finally, the issue of applying a pseudo-steady state assumption to unsteady-state

fermentation data requires more careful consideration, since the ﬂuxes of NADH2 (and

potentially of all the intermediates) in the cell would be expected to change during

179

oscillations in CO consumption. The addition of another piece of data such as CO

consumption could alleviate this problem by eliminating the need for reliance on the
NADHz balance.

CHAPTER 6

CONCLUSIONS AND RECOMMENDATIONS

The research contained in this dissertation dealt with biological conversion of CO gas to
acids and alcohols, using Butyribacterium methylotrophicum. The investigation was
focused in three major areas: metabolic regulation and tolerance, cell recycle fermentation,
and metabolic modeling. Several conclusions and observations were made during the
course of this work, as discussed below.

Batch studies were conducted to determine the effect of H2S on the growth and
metabolism of B. methylotrophicum. H2S concentrations in the range of those present in
typical synthesis gas mixtures were examined. Growth rates and product compositions
were unaffected by total liquid phase sulﬁde concentrations ranging from 0.0 to 2.0
mmol/L, which corresponds to an H2S partial pressure of approximately 04.4%. An
analysis of the effect of H28 equilibrium partitioning, media pH, and gas-liquid
equilibration rates on the experimental system was also conducted.

B. methylotrophicum can convert CO or hydrogen and carbon dioxide to acetic acid.
Studies have shown that the CO metabolism of this organism can be shifted in favor of
butyric acid and ethanol and butanol by lowering the fermentation pH. In this work,
growing cells of B. methylotrophicum were incubated with CO gas under varying
conditions of external alcohol concentration, acid concentration, and broth pH.
Concentrations of ethanol and butanol above 3.0 g/L were found to be toxic to the cells.
However, preliminary stationary phase experiements showed no inhibition at
concentrations of 5 g/L and 7 g/L for ethanol and butanol, respectively. Under conditions
of high pH and high acetate concentration (>8 g/L), B. methylotrophicum exhibited
consumption of both CO and acetate, with net production of butyrate, butanol, ethanol and

cell mass. Similarly, under conditions of high pH and moderate butyrate concentration (>2

180

181

g/L), B. methylotrophicum exhibited consumption of both CO and butyrate, with net
production of acetate, butanol, ethanol and cell mass. These studies indicate that regulation
of CO metabolism in B. methylotrophicum can be achieved in both the growth and
stationary phases using a variety of environmental conditions.

Fermentation studies were conducted using a cell recycle system integrated with a
chemostat to increase cell density and bias the cell population towards the stationary
(resting) state. Results indicate three metabolic regimes for this system. Between pH
values of 6.0 and 7.2, steady-state operation was achieved, and a 5-10 fold increase in
acid and alcohol production was observed relative to chemostat operation without cell
recycle. At pH values of 5.75 and 5.5, oscillations in butyrate and acetate production were
observed, with relatively steady alcohol production. Importantly, butanol was the major
product over a 72 hour period during initial start-up of the continuous cell recycle system.
At pH values of 5.35 and 5.1, the cells exhibited a death phase response. These results
have demonstrated that during high cell density, high productivity CO fermentation,
fermentation pH induces speciﬁc metabolic responses. Furthermore, these responses give
insight to the mechanisms controlling product formation, speciﬁcally 4-carbon products,
directly from C0 gas.

A metabolic model was developed to analyze the cell recycle fermentation data. This
comprised a system of linear equations based on known metabolic pathways for CO
metabolism in B. methylotrophicum. The model incorporated a hypothetical reaction for
generation of ATP via electron-transport phosphorylation, as well as known substrate-
level phosphorylation routes for ATP synthesis. The theoretical molar yields for ATP were
shown to be a function of fermentation pH, and to be strongly dependent on the rate of cell
mass production. ETP was concluded to be a required portion of the overall energetics,
based on the energetic analysis of steady-state results. A preliminary analysis of unsteady-
state results suggested that further, as yet unknown ATP generating pathways, either ETP-

based or via some other mechanism, must be present in B. methylotrophicum. Oscillations

182

in net theoretical ATP yields were shown to be concommitant with oscillations in acetate
and butyrate formation, indicating that the net ATP charge could play a regulatory role in
directing carbon and electron ﬂow dming CO metabolism.

The mechanism of the pH effect, where low pH values inhibit acetate formation and cell
growth, is complicated and not fully understood. A case for mass-action effects, or
feedback inhibition of acetate formation, was presented based on known regulation
mechanisms in other organisms. This mechanism presumes the regulatory species is the
protonated form of the acid, and thus regulation of metabolism is a function of both
concentration and of pH. Since formation of acetate is the main reaction contributing to
ATP synthesis via SLP, the regulatory effects of pH and of ATP charge, if present, would
be interdependent. Although the data support this theory, conﬁrmation of the several
hypotheses requires a biochemical analysis of both carbon and electron ﬂow and of the
cellular energetics.

Experimentally, several improvements in the present system are needed for further
development, both of which center on data aquisition. By far, the major limitation with the
present work is the reliance on carbon and electron balances for calculation of key

fermentation parameters such as cell mass production, CO uptake, and C02 production.
The use of an inert tie-component in the gas substrate, to allow calculation of CO and C02
gas partial pressures, is a potential solution to this issue. One study that is of paramount
importance is a parametric sensitivity analysis on the fermentation equation for cell mass
production used in the metabolic model. Small variations in the magnitudes of the
coefﬁcients for electron and ATP requirements can potentially result in large differences
for calculated reaction rates, based on the electron balance for NADH. The equation for
cell mass production should thus be closely examined and its accuracy conﬁrmed.

The success of the cell recycle fermentations in increasing fermentation productivity,
roughly 8-10 fold, and of elucidating a capacity for primary production of butanol from
C0 gas, has greatly enhanced the potential of the synthesis gas bioconversion concept.

183

However, there are still several important issues which must be addressed for this system,
most of which deal with the reaction rate for CO uptake. Other organisms have exhibited
much faster rates of growth and product formation from CO, but produced only acetate
and ethanol. A strong potential exists for ﬁnding an organism capable of high CO uptake
rates as well as formation of 4-carbon products such as butyrate and butanol. Thus, an
isolation effort directed at finding new and improved microorgansims may be the most
promising direction for future work.

Another area of research with some promise is in the use of cosubstrates such as
methanol during CO metabolism, or perhaps just methanol alone. These approaches would
facilitate increases in reaction rates, and are not outside the realm of synthesis gas
feedstock utilization. A speciﬁc methanol to butanol fermentation would perhaps be the
goal of such work, since acetate from methanol and CO is already well established
commercially. B. methylotrophicum has exhibited the versatility for such approaches.
However, whether with B. methylotrophicum or some other organism, the ultimate
potential for biological conversion of synthesis gas will require further development for

such processes to become a commercial reality, in the face of cm'rent chemical technology.

APPENDIX

APPENDIX

Review of the Carbon and Electron Balancing Method

The method of carbon and electron balancing for determination of biological
fermentation stoichiometries is a powerful tool when data for consumption or production
of one or two species is unavailable. These balances also allow calculation of carbon and
electron recovery in the fermentation products. Carbon balancing, as detailed below, is
simply a material balance on the reaction, based on the moles of carbon atoms reacted.
Electron balancing is a balance on the available electron equivalents and is analogous to an
energy balance on the fermentation reaction.

Speciﬁcally, electron balances are derived from molecular energy changes associated
with changes in oxidation number due to electron transfer between species. The atomic
oxidation number is characteristic of the valence electron(s) of the atom, which
stoichiometrically guide the reaction and which for carbon containing compounds can
assume several values. The energy content of carbohydrate molecules is related to the
oxidation number, with highly reduced compounds being energy-rich. Changes in
oxidation number arise from electron transfer between species, often coupled to redox
reactions involving electron carriers such as NADH.

Within this setting, the reductance degree of a compound is deﬁned as the number of
electrons available for transfer in oxidative catabolism, and is established on a basis of one
gram atom carbon. For atoms, several reductance degrees are,

C: +4 H: +1 N: -3 O: -2
and thus for several compounds can be calculated as
CO: +2 CH30H: +6 CH3COOH: +8 C02: 0

184

185

and ad inﬁnitum. The reductance degree therefore forms a basis for energy balancing
biological reactions where the total electrons supplied by the reactants, or substrates, must
equal the total electrons obtained by the products. The material balance for carbon is also
achieved in this fashion.

As a speciﬁc example related to calculations made during this study, the carbon and
electron balances for a 100% CO batch fermentation will be illustrated. Fermentation data
include product concentrations of acetic acid, butyric acid, and cell mass, but no data is
obtained for CO consumption and C02 production. For this anaerobic reaction, a typical
stoichiometry is of the form

aCO + bH20 + cNH3 --> dCOz + eCH3COOH + fC3H7COOH + gCells
where the lower case letters represent total moles of each species, with the cell mass term
expressed with a molecular formula based on one gram-mole of carbon. If this general
reaction is balanced with respect to available electron contents, with the Cells term having a
previously derived value of 4.21 for B. methylotrophicum, the resulting equation is

a(2) + b(0) + c(0) ---> d(0) + e(8) + f(20) + g(4.21)
which is equivalent to
a(2) ---> e(8) + f(20) + g(4.21).
The product concentration data for acetate, butyrate, and cell mass is then converted to
total moles and inserted in this equation, where
a(2) --> 010(8) + 0.0031(20) + 0.015(4.21)
and the value for a, the total moles CO consumed, calculated. This value is
a = 0.4624 moles CO consumed.

If the reaction is then balanced with respect to carbon, the equation is of the form

aCO ---> dC02 + eCH3COOH + fC3H7COOH + gCells
which upon insertion of the total carbon atoms in each molecular formula yields

a(l) ---> d(1) + e(2) + f(4) + g(l).

186

The product data is then inserted and the value of b, the moles C02 produced, calculated
from
0.4624(1) ---> d(l) «1» 010(2) + 0.0031(4) + 0.015(1)

which yields
b = 0.2350 moles C02 consumed.

The overall carbon and electron balance for the fermentation can now be normalized to
an arbitrary amount of substrate, in this case 4 moles of CO, which results in a ﬁnal

fermentation balance of

4CO ---> 2.03C02 + 0.86CH3COOH + 0.027C3H7COOH + 0.13Cells

which is a representative balance derived for 100% CO batch fermentation.

LIST OF REFERENCES

LIST OF REFERENCES

Abbanat, D. R. and J. G. Ferry. 1990. Synthesis of Acetyl Coenzyme A by Carbon
Monoxide Dehydrogenase Complex from Acetate-Grown Methanosarcina thermophila.
J. Bacterial. 172:7145-7150.

Agreda, v. H. 1988. Acetic Anhydride from Coal. Chemtech. 11111-250253.

Baer, S. H., H. P. Blaschek, and T. L. Smith. 1987. Effect of Butanol Challenge and
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