ANALISIS COMPARATIVO DE USOS DE GLICEROL

Design and Analysis of Technological Schemes for Glycerol Conversion to Added Value Products

John Alexander Posada Duque

Universidad Nacional de Colombia

Facultad de Ingeniería y Arquitectura, Departamento de Ingeniería Eléctrica, Electrónica y

Computación

Manizales, Colombia

2011

Diseño y Evaluación de los Esquemas Tecnológicas para la

Conversión de Glicerol en Produtos de Valor Agregado


Universidad Nacional de Colombia Sede Manizales


Computación

Manizales, Colombia

Design and Analysis of Technological Schemes for Glycerol Conversion to Added Value Products


Thesis submitted in partial fulfillment of the requirements for the degree of:

Doctor of Philosophy in Engineering

Advisor:

Ph.D., M.Sc, Chemical Engineer Carlos Ariel Cardona Alzate

Research line:

Chemical and Biotechnological Process Engineering

Research group:

Chemical, Catalytic and Biotechnological Process

Universidad Nacional de Colombia


Computación

Manizales, Colombia

2011

______________________

A mi madre y mi hermana por su apoyo,

a Patricia por su valiosa presencia.

Acknowledgment I would like to express thanks to God and to my mother for the most essential reason, the

life. I also would like to express sincere thanks to my advisor, Dr. Carlos Ariel Cardona

Alzate, who has been for more than six years the most important influence not only in my

career but also in my life. Kind thanks to Dr. Ramon Gonzalez for receiving me at his

laboratory during the internship. Special thanks to my research fellows, Luis Rincon,

Julian Quintero, and Javier Naranjo, for their friendship and unconditional help. And

thanks to Patricia Arevalo and to my friends.

Thanks to the National University of Colombia for the financial support, to the Research

Office of National University of Colombia branch Manizales for the financial support in the

internship in United States, to the Research and Extension Projects Office of National

University of Colombia branch Manizales for the financial support in air tickets to United

States, to the Department of Electricity, Electronic and Computational Engineering of

National University of Colombia branch Manizales for the financial support for attending to

congresses.

Resumen y Abstract V

Resumen Uno de los principales problemas relacionados con la creciente industria del biodiesel es

la sobreproducción de glicerol, que se obtiene en una relación en peso de 1/10 (glicerol/

biodiesel). Lo que ha llevado a que su precio de venta caiga en un orden de magnitud.

Así, la gran cantidad de glicerol co-producido puede ser usada como una materia prima

renovable y de bajo costo para producir compuestos químicos y combustibles. Aquí se

analiza la conversión química y bioquímica de glicerol hacia productos de valor agregado

basado en criterios tecno-económicos. Entonces se consideraron nueve productos

finales (vía química: gas de síntesis, acroleína y 1,2-propanodiol; vía bioquímica: etanol,

1,3-propanodiol, ácido D-láctico, ácido succínico, ácido propiónico y poly-3-hidroxibuti-

rato). Además, un total de 27 esquemas tecnológicos fueron diseñados, simulados y

evaluados económicamente utilizando Aspen Plus y Aspen Icarus Process Evaluator.

Como una conclusión, una plataforma de biorefinerias basada en glicerol fue obtenida

para la producción rentable de combustibles fósiles y bioclásticos.

Palabras clave: Conversión de glicerol, diseño de procesos, simulación de procesos,

evaluación de procesos, biorefinerias basadas en glicerol.

VI Glycerol Conversion to Added Value Products

Abstract An important concern related to the growing biodiesel industry is the over-production of

raw glycerol as by-product, which is obtained in a weight ratio of 1/10 (glycerol/biodiesel).

This fact had led to a 10-field drop of its sale price. Thus, the large amount of by-

produced glycerol can be used as low-cost and renewable feedstock in order to produce

chemicals and fuels. Here, the chemical and bio-chemical conversion of glycerol to

added-value products was analyzed based on techno-economic criteria. In this way, nine

final products (for chemical conversion: syn-gas, acrolein, and 1,2-propanediol; while for

fermentative conversion: ethanol, 1,3-propanediol, D-lactic acid, succinic acid, propionic

acid, and poly-3-hydroxybutyrate) were considered. And a total 27 technological schemes

were designed, simulated, and economically assessed, using Aspen Plus and Aspen

Icarus Process Evaluator. As a conclusion a glycerol-based platform for biorefineries was

obtained for the profitable production of fuels, chemicals, and bio-plastics.

Keywords: Glycerol conversion; process design; process simulation; process

assessment; glycerol-based biorefineries.

Contents VII

Table of Contents

ACKNOWLEDGEMENTS IV Pág.

RESEMEN V ABSTRACT VI TABLE OF CONTENTS VII LIST OF FIGURES X LIST OF TABLES XIII 1. Introduction 1

1.1. Application field and Motivation 2 1.2. Thesis’ objectives 5 1.3. Thesis’ structure 5

References 7

2. Chapter 2: The glycerol’s world 9 2.1. Overview 9 2.2. Biodiesel industry 10 2.3. Glycerol market and its oversupply problem 12 2.4. Glycerol as raw material 13

References 14

3. Chapter 3: Methodology for processes design and analysis 17 3.1. Processes design 17 3.2. Processes simulation 20 3.3. Processes assessment 22

References 24

4. Chapter 4: Separation and purification of glycerol 25 4.1. Commercial qualities of glycerol 25 4.2. Effect of the feedstock for biodiesel production on glycerol composition 27 4.3. Conventional purification process 28 4.4. Alternative purification process 29

VIII Glycerol Conversion to Added Value Products

4.5. Simulation of the glycerol purification process 31 4.6. Economical assessment for glycerol purification processes 33 4.7. Conclusions 34

References 35

5. Chapter 5: Chemical conversion of glycerol 36

5.1. Oxidation 36 5.2. Reduction 39 5.3. Etherification 41 5.4. Pirolysis and gasification 41

References 43

6. Chapter 6: Biochemical conversion of glycerol 47 6.1. 1,3-Propanediol 47 6.2. Ethanol 48 6.3. Poly-3-hydroxybutirate 51 6.4. D-Lactic acid 53 6.5. Succinic acid 54 6.6. Propionic acid 58

References 59

7. Chapter 7: Study cases for chemical conversion of glycerol 69 7.1. Generalities 69 7.2. Acrolein production 70 7.3. Hydrogen production 73 7.4. 1,2-propanediol production 76 7.5. Economic assessment 78 7.6. Conclusions 81

References 81

8. Chapter8: Study cases for biochemical conversion of glycerol 85 8.1. 1,3-Propanediol production 85 8.2. Ethanol production 104 8.3. PHB production 115 8.4. D-Lactic acid production 124 8.5. Succinic acid production 138 8.6. Propionic acid production 148 8.7. Economic assessment 153 8.8. Conclusions 167

References 169

Contents IX

9. Chapter 9: Experimental setup for glycerol fermentation to PHB 183 9.1. Generalities 183 9.2. Materials and methods 186 9.3. Analytical Methods 187 9.4. Results and discussion 189

References 192

10. Chapter 10: Conclusions 194

11. Chapter 11: List of Publications 199

Figures X

List of Figures Page

Figure 1.1. World biodiesel production and capacity 3

Figure 1.2. Global biodiesel production by feedstock 4

Figure 3.1. Hierarchical decomposition according to the "onion diagram" 19

Figure 3.2. The process design method based on the so called breadth-first 20

Figure 4.1. Flowsheet of conventional schemes for glycerol purification 29

Figure 4.2. Flowsheet of the Ambersep BD50 process 30

Figure 4.3. Simplified flowsheet for raw glycerol purification 31

Figure 5.1. Possible products for glycerol oxidation 38

Figure 6.1. Schematic representation of glycerol degradation process on the part 49 of Escherichia coli, on non fermentative process.

Figure 6.2. Main metabolic pathways for fermentative degradation of glycerol 50 by Escherichia coli.

Figure 6.3. Pathways involved in the microaerobic utilization of glycerol in E. coli 54

Figure 6.4. Products that can be synthesized from succinic acid 55

Figure 6.5. Pathways involved in the micro aerobic utilization of glycerol 57

Figure 7.1. Simplified flowsheet for acrolein production by glycerol dehydration 71

Figure 7.2. Simplified flowsheet for hydrogen production by gasification 74

Figure 7.3. Simplified flowsheet for 1,2-propanediol production by hydrogenolysis 77

Figure 8.1. Hysteresis loops and multiple steady states 89

Figure 8.2.a) 1,3-propanediol volumetric productivity, (the column in the right side 90 gives the scale). b) Region of multiplicity of steady states, optimal productivity for each dilution rate, global optimal productivity, and wash-out line.

Figures XI

Figure 8.3. 1,3-Propanediol productivity and concentration in the second 91 fermentation stage

Figure 8.4. Volumetric productivity in the second fermentation stage using the 92 optimal dilution rate obtained by the model 1 for the first fermentation stage, (the column in the right side gives the scale).

Figure 8.5. Product of productivities of both fermentation stages using the optimal 93 dilution rate obtained by the model 1 for the first fermentation stage, (the column in the right side gives the scale).

Figure 8.6. Acetylation reaction of 1,3-propanediol with iso-butyl aldehyde to 95 2-iso-propyl-1,3-dioxane

Figure 8.7. Simplified flowsheet for 1,3-propanediol production from raw glycerol 98

Figure 8.8. Residue map curves for the reactive system 99

Figure 8.9. Direct separation with fed 0.377645/0.622355–2iP13DO/Water 100

Figure 8.10. P/W ratio, Direct Separation (XF: 0.377645/0.622355-2iP13DO/water) 101

Figure 8.11. iso-Volatility curve (Water–iso-Butyraldehyde–2-iso-Propil-1,3-Dioxane) 101

Figure 8.12. Simplified flowsheet of fuel ethanol production from glycerol at 104 88 wt % and 98 wt %.

Figure 8.13. Stages for ethanol production from sugar cane, corn, and crude 107 glycerol

Figure 8.14. Simplified flowsheet for ethanol production from: (A) sugar Cane 108 and (B) Corn

Figure 8.15. Flowsheet for the integrated process of combined biodiesel and 112 bioethanol production

Figure 8.16. Flowsheets for PHB production from glycerol (88 or 98 wt %) 120

Figure 8.17. Scheme for the simulation procedure to synthesize PHB from crude 124 glycerol

Figure 8.18. Complex formed during the reactive extraction process of 133 D-lactic acid

Figure 8.19. Simplified flowsheet for D-lactic acid production from raw glycerol 134

Figure 8.20. Reaction complexes of succinic acid, formic acid and acetic acid 144 with TOA

Figure 8.21. Simplified flowsheet for succinic acid production from raw glycerol 144

XII Glycerol Conversion to Added Value Products

Figure 8.22. Simplified flowsheet for propionic acid production from raw glycerol 150

Figure 9.1. Accumulation profile I of Bacillus megaterium 189

Figure 9.2. Accumulation profile II of Bacillus megaterium 190

Figure 9.3. Accumulation profile III of Bacillus megaterium 190

Figure 9.4. Accumulation profile IV of Bacillus megaterium 191

Tables XIII

List of Tables Page

Table 3.1. Used costs and prices for the economic assessment 23

Table 4.1. Quality specifications for the main qualities of glycerol 25

Table 4.2. Fatty acid profile of vegetable and used oils 27

Table 4.3. Composition of the glycerol layer obtained by decantation during 28 the biodiesel production from different feedstocks

Table 4.4. Simulation results for raw glycerol purification process 32

Table 4.5. Purification costs (PC) of raw glycerol (US$/L) 33

Table 7.1. Simulation results for dehydration process from glycerol 72

Table 7.2. Simulation results for gasification process from glycerol 74

Table 7.3. Simulation results for hydrogenolysis process from glycerol 78

Table 7.4. Production costs for glycerol conversion to added-value 79

Table 7.5. Percentage of Production costs for glycerol conversion to added-value 79

Table 8.1. Results summary for each optimization model 94

Table 8.2. Fermentation results for the three considered scenarios 96

Table 8.3. stoichiometric reactions for each scenario and each fermentation stage 97

Table 8.4. Singular Points ** - Acetilation System of 1,3-PD* with 2iP13DO* 99

Table 8.5. Summary of the main simulation results for 1,3-propanediol production 102 from glycerol

Table 8.6. Data representing the behavior of the downstream process 104

Table 8.7. Simulation results for fuel ethanol production from glycerol 106

Table 8.8. Main input data and operation conditions used in the simulation process 113

XIV Glycerol Conversion to Added Value Products

Table 8.9. Main process streams for ethanol production from lignocellulosic biomass 114

Table 8.10. Process conditions for glycerol fermentation 116

Table 8.11. PHB Extraction Methods 118

Table 8.12. Process conditions for PHB recovery: Downstream Process I 121

Table 8.13. Process conditions for PHB recovery: Downstream Process II 121

Table 8.14. Process conditions for PHB recovery: Process III 122

Table 8.15. Downstream processes for lactic acid recovery from a fermentation broth 131

Table 8.16. Base information for the glycerol fermentation to D-lactic acid 132

Table 8.17. Stoichiometry for glycerol fermentation to D-Lactic Acid by Engineered 134 E. coli

Table 8.18. Summary of the main simulation results for D-lactic acid production 135 process

Table 8.19. Data representing the behavior of the downstream process for D-lactic 137 acid production Table 8.20. Base information for the glycerol fermentation to succinic acid 141

Table 8.21. Stoichiometry for glycerol fermentation to succinic acid by Engineered 141 E. coli

Table 8.22. Removal efficiency (%) of the carboxylic acids from the fermentation 142 broth.

Table 8.23. Summary of the main simulation results for succinic acid production 145 process

Table 8.24. Data representing the behavior of the downstream process for succinic 147 acid production

Table 8.25. Stoichiometry of the fermentation process for each scenario 149

Table 8.26. Summary of the main simulation results for prpionic acid production 151 process.

Table 8.27. Data representing the behavior of the downstream process for propionic 153 acid production

Table 8.28. Economic results for raw glycerol conversion to 1,3-propanediol: Cost 154 (USD$/kg) and Share (%)

Table 8.29. Bioconversion costs (BCCs) for fuel ethanol production form raw glycerol 166

Tables XV

Table 8.30. Global production costs (GPCs) for fuel ethanol production from 157 raw glycerol

Table 8.31. Discriminated costs for integrated biodiesel and raw-ethanol production 158 from oil palm

Table 8.32. Total PHB production costs from crude glycerol through raw glycerol 159 (88 wt %) and pure glycerol (98 wt %).

Table 8.33. Main producers of PHA in the world 160

Table 8.34. Economic results for raw glycerol conversion to D-lactic acid: Cost 162 (USD$/kg) and Share (%)

Table 8.35. Economic results for raw glycerol conversion to succinic acid: Cost 164 (USD$/kg) and Share (%)

Table 8.36. Economic results for raw glycerol conversion to propionic acid: Cost 166 (USD$/kg) and Share (%).

Table 9.1. Some microorganisms PHB producer from different agroindustrial wastes 185

Table 9.2. Comparison between experimental results for glucose and glycerol 191 fermentation to PHB

1. Introduction

Fossil sources have diminished significantly because they have been used as the main

raw material for the current economy and life style. For instance, large scale products

such as transportation fuels and daily use components are obtained from petrochemical

industry. Furthermore, the increasing demand of fossil sources from developing

economies (like China and India) and speculations about oil reserves availability have

caused high crude oil prices. In fact, experts predict the end of cheap oil in 2040 at the

latest [1], which are currently and again above USD$ 100 per barrel. This economic issue

added to the environmental conscience, which is focus on the problems derived from

pollution and accumulation of greenhouse gases, have taken to develop alternative

technologies in order to produce sustainable fuels and chemicals using renewable

resources. In this way, biodiesel and bioethanol are the most important technological

platforms for liquid fuels production.

Although biofuels such as biodiesel and bioethanol represent a renewable, convenient,

and environmental friendly alternative for fossil fuels substitution, they also cause

concerns in relation to their economic viability. Implementation of biorefineries as an

additional process to the biofuels production is an interesting alternative to both overcome

the limited profitability of these technologies and use the generated sub-products.

Therefore, the concept of biorefinery could be especially advantageous if the conversion

of by-products or wastes to added-value products is considered [2].

Glycerol as the main by-product on biodiesel production is obtained at high concentration

in a weight ratio of 1/10 (glycerol/biodiesel). Moreover, the growing market of biodiesel

has generated a glycerol oversupply, where its production increased 400% in a two years

period and consequently the glycerol commercial price fell down near to 10 fold during the

same period of time [3]. As a result of the low prices of glycerol, traditional producers such

as Dow Chemical, and Procter and Gamble Chemicals, stopped its production [4].

2 Glycerol Conversion to Added Value Products

Since glycerol sales have represented an important profitability for biodiesel industry, it is

reasonable that low prices of glycerol could impact the economy of biodiesel producers

negatively. For that reason, the correct exploitation of glycerol as raw material should be

focused on its transformation to added-value products. Thus, the use of glycerol is a high-

priority topic for managers and researchers related to biofuels production. In this sense,

the establishment of glycerol’s biorefineries able to co-generate added-value products is

an excellent opportunity not only to raise the profitability but also to produce other

chemicals from a biobased raw material.

1.1 Application field and motivation

In order to analyze the glycerol conversion possibilities, this highly functional molecule

has been identified as a potential raw material for organic synthesis of many

intermediates and chemical products. Chemically glycerol can be transformed by many

ways such as oxidation, hydrogenolysis, etherification, pyrolysis, and gasification. Thus,

different kinds of products such as acrolein, 1,2-propanediol, polyglycerols, syn gas,

among many others compounds can be chemically obtained. On the other hand, because

of glycerol is a structural component of many lipids, it can also be biochemically

transformed to added value compounds. Some products of glycerol fermentation are: 1,3-

propanediol, ethanol, propionic acid, citric acid, lactic acid, poly-3-hydroxybutirate, and

biosurfactants. Then, due to the wide variety of potential products from glycerol, its

biorefineries are an excellent commercial opportunity. In this way, it is necessary to

determine the most appropriate alternative for glycerol transformation. In this study, the

process design and the assessment of different technological schemes for glycerol

transformation to added-value components is systematically performed considering

technologic and economic indicators.

Biodiesel production sector is a dynamic industry with a rapid global market growth. For

instance, over the past decade the biodiesel production was governmentally driven aiming

to the development of large scale industries. Thus, Europe took the lead with more than

1.6 mill Tons of biodiesel produced in 2002 (at capacities of approx. 2.1 mill Tons), while

in the USA approx. 40.000 T were produced [5]. Furthermore, in 2008 the global biodiesel

production reached more than 11.1 mill Tons (see Figure 1.1), representing around 1 % of

all diesel consumption of the USA and between 2-3 % of the total transportation

1. Introduction 3

consumption in Europe [6]. Even though, Europe represents 80% of global biodiesel

production and consumption, the U.S. is increasing its production at a faster rate than

Europe, while Brazil is expected to surpass the U.S. and European biodiesel production

by the year 2015 [6].

Figure 1.1. World biodiesel production and capacity.

Currently, new economic and environmental concerns are leading to create governmental

incentives targeting to a combination of: reduction of petroleum imports and increase of

production and consumption of renewable fuels. In this way, Europe, Brazil, China, and

India each have targets to replace 5% to 20% of total diesel with biodiesel [6]. In addition,

if governments promote the development of second generation biofuels (and their

production using alternative and non-food feedstocks) throughout investment and politics,

the prospects for biodiesel market will be early reached. Figure 1.2 shows the expected

biodiesel production from different feedstocks where the share in total biodiesel

production from edible vegetable oil could decrease from almost 90% to about 75% by

2019. This expected change is due to the development of biodiesel production from

jatropha mainly in India and to the increasing use of animal fats to produce biodiesel in

the USA. Also, biomass based biodiesel could represent almost 6.5% of total biodiesel

production by 2019.

0

5

10

15

20

25

30

35

2001 2002 2003 2004 2005 2006 2007 2008 2009

Production Capacity


Figure 1.2. Global biodiesel production by feedstock

Under the above described situation, high quantities of raw glycerol are continuously

produced since it is obtained in ratio of 9 wt % respect to the produced biodiesel. Then, in

order to avoid both economic and environmental drawbacks related to the use and

disposal of glycerol, new applications for glycerol must be proposed. Even though, the

most traditional applications of glycerol have been related to its use as additive in: food,

tobacco, pharmaceuticals and medicine, and for the synthesis of trinitroglycerine, alkidic

resins, and polyurethanes, one of the most attractive alternatives for glycerol utilization is

as feedstock for producing added-value compounds such as: bioplastics, platform

chemicals, and fuels. Thus, because of both the low prices and high availability of

glycerol, this compound could be a great opportunity to make money through biorefineries

built adjacent to the biodiesel production plant.

From a chemical view of point, glycerol is a highly versatile molecule with two primary

hydroxyl groups and a secondary hydroxyl group which offers different reaction

possibilities. Meanwhile from a biochemical view of point, the glycerol molecule is

abundant in nature in the form of triglycerides (a chemical combination of glycerol and

fatty acids) which are the major constituents of nearly all vegetable oils and animal fats.

Thus, the high functionality and occurrence in nature of glycerol allow it to be transformed

by a chemical route or a fermentative way, as it was above indicated.

1. Introduction 5

In this way, the most important possibilities for glycerol transformation to added-value

compounds are here reviewed and methodologically assessed by mean of processes

engineering tools such as: process design, process simulation, and economic evaluation.

And finally all the analyzed possibilities are systematically compared.

1.2 Thesis’ objectives

This thesis aims to design and assess technological schemes for the conversion of raw

glycerol obtained during the biodiesel production to added-value products, in order to

identify the best alternatives from a technical and economic view of point. Thus, this

research required: (i) to identify and select the most promissory possibilities for glycerol

transformation, (ii) to simulate and assess the chosen technological schemes and

scenarios for the several identified potential products from glycerol, and (iii) to compare

these technological schemes based on economic criteria.

1.3 Thesis’ structure This thesis presents the results of different studies that have been already published or

are under review for their publication.

The thesis is accordingly divided into the following

chapters:

- Chapter 2: The glycerol’s world

This chapter introduces to the reader with the current status of glycerol as the by-product

on biodiesel production and discusses the glycerol problem related to its oversupply.

Additionally, the main uses of glycerol as additive and its market are also presented.

Finally, the glycerol conversion possibilities are described.

- Chapter 3: Methodology to design and analyze processes based on simulation tools.

This chapter details the used methodology for the process design, processes simulation,

and process assessment in both cases, i.e., chemical and fermentative conversion of

glycerol.


- Chapter 4: Separation and purification of glycerol

This chapter presents the requirements for the most important commercial qualities of

glycerol, as well as the influence of the feedstock used for biodiesel production on the

glycerol layer. Additionally, conventional and non-conventional processes for raw glycerol

purification are discussed. Finally, the purification costs of raw glycerol up to different

commercial qualities are obtained based on simulation and economic assessment tools.

- Chapter 5: Chemical conversion of glycerol

This chapter reviews the alternatives for chemical conversion of glycerol by different

reaction ways such as: oxidation, reduction (hydrogenolysis), etherification, pirolysis, and

gasification. Conversion levels, yields, selectivities, and productivities are also presented.

- Chapter 6: Biochemical conversion of glycerol

This chapter reviews the alternatives for fermentative conversion of glycerol by different

strains. Fermentation products such as: 1,3-propanediol, ethanol, lactic acid, succinic

acid, propionic acid, poly-3-hydroxybutyrate, and biosurfactants are discussed.

Additionally, conversion levels, yields, selectivities, and productivities are presented.

- Chapter 7: Cases of study for chemical conversion of glycerol

This chapter presents the flowsheets, simulation results, and economic assessments for

the chemical conversion of glycerol to: acrolein, 1,2-propanediol, and hydrogen.

- Chapter 8: Cases of study for biochemical conversion of glycerol

This chapter presents the flowsheets, simulation results, and economic assessments for

the fermentative conversion of glycerol to: 1,3-propanediol, ethanol, D-lactic acid, succinic

acid, propionic acid, and poly-3-hydroxybutyrate.

- Chapter 9: Experimental setup

This chapter shows the experimental setup performed for poly-3-hydroxybutyrate

production from glycerol using two strains: cupriavidus necator and bacillus megaterium.

1. Introduction 7

- Chapter 10: Conclusions

This chapter contains the general conclusions of the thesis and also presents the

contributions made during this work. Finally, some recommendations for future works are

given.

- Chapter 11: List of publications and submitted papers

This chapter shows the published results throughout scientific meeting, papers, book

chapters, invited book chapters, and books. Also, a list containing the submitted papers

was included.

References

[1] Posada J.A., Orrego C.E., Cardona C.A. 2009. Biodiesel production: Biotechnological

approach. International Review of Chemical Engineering (I.Re.Che.), 1(6):571-580.

[2] Yazdani S.S. and Gonzalez R., 2007. Anaerobic fermentation of glycerol: A path to

economic viability for the biofuels industry. Current Opinion in Biotechnology, 18:213–219

[3] Posada J.A., Cardona C.A., 2010. Análisis de la refinación de glicerina obtenida como

co-producto en la producción de biodiesel (Validation of glycerin refining obtained as a

by-product of biodiesel production). Ingeniería y Universidad 14:2-27.

[4] Posada J.A., Cardona C.A., Cetina D.M., Orrego C.E., 2009. Bioglicerol como materia

prima para la obtención de productos de valor agregado (Bioglycerol as raw material to

obtain added value products). En: Cardona C.A. (ed). Avances investigativos en la

producción de Biocombustibles (Reasearching advances for biofuels production).

Manizales: Ed. Universidad Nacional de Colombia sede Manizales. p. 103-127. ISBN:

978-958-44-5261-0.

[5] Worldwide review on biodiesel production. Austrian Biofuels Institute,

www.biodiesel.at, 2003.

[6] Biodiesel 2020: Global Market Survey, Feedstock Trends and Market Forecasts.

http://www.biodiesel.at/�

2. The Glycerol’s World

This chapter describes the relationship between the market of both glycerol and biodiesel,

and it also discusses the influence of the growing biodiesel production on the commercial

prices of glycerol. Additionally, the potential of raw glycerol for biorefineries developing

using it as a main feedstock is presented. Finally, an overview on the possibilities of

glycerol transformation by chemical and biochemical routes is given.

2.1 Overview

The glycerol molecule (1,2,3-propanetriol) is a highly reactive tri-alcohol which has two

primary and a secondary hydroxyl groups. Some physical characteristics are: water

soluble, colorless, odorless, viscous, and hygroscopic; with a specific gravity of 1.261 g

mL-1, melting temperature of 18.2 °C, and a boiling temperature of 290 °C (accompanied

by decomposition). Chemically, glycerol is able for reacting with a stable alcohol under

most operational conditions, and it is basically non-toxic to human health and

environment. The key of its usefulness is the particular combination among its

physicochemical properties, compatibility with other substances, and easy handling. Due

to these particular properties set, glycerol has found more than 1500 end-uses or large

volume applications.

Glycerol is a commodity chemical obtained mainly as by-product in the oleochemical and

biodiesel industry; meanwhile glycerin is the commercial name for mixtures containing

high quantities of glycerol. This molecule is one of the most versatile substances known

due to its unique combination of physical and chemical properties, which allows it to be

used in multitude of products and additionally it is often used as: humectant, plasticizer,

emollient, thickener, solvent, dispersing medium, lubricant, sweetener, and antifreeze.

Glycerol is naturally combined with triglycerides in all animal fats and vegetable oils,

representing about 10% of these materials. This component is derived from fats and oils


during fatty acids and soap production, or by the transesterification process with alcohols

for biodiesel synthesis. Although glycerol can also be produced synthetically by

petrochemical processes from epichlorohydrin and using propylene as raw material, such

processes are no longer conducted at the industrial level [1].

2.2 Biodiesel industry

The glycerol’s world has a complex behavior since it is by-produced with biodiesel, and in

addition its price is related to the no-predictable network of both its supply and demand.

This is a typical behavior for a commodity used as additive in many applications and now

being used as raw material for the production of platform chemicals, bioplastics, and

biofuels. Here the most important topics related to the glycerol industry are elucidated.

Biodiesel is defined as a clean burning fuel used for diesel engines, manufactured from

renewable sources (vegetable oils, animal fats, or used cooking oils) and short chain

alcohols (methanol, ethanol, or butanol), to produce a methyl, ethyl or butyl esters fatty

acids mixture. A vegetable oil usually contains up to 14 different kinds of fatty acids [2].

Most of biodiesel production processes were developed in the early 40’s, during World

War II by explosives manufacturers searching for a simpler way to obtain glycerol. Now,

biodiesel is commercially produced from agricultural products such as rapeseed, soy

bean, and palm oils. Also, other high fatty acid feedstocks such as: used frying oil, grease

trap waste oil, and waste tallow or lard, have been used. Several variables as local

availability, cost, government support, and fuel performance must be analyzed in order to

choose the best feedstock, since biodiesel production costs are highly dependent upon

the feedstocks price.

Biodiesel is a fuel with low viscosity and pour point, non-toxic, and biodegradable, which

is also cleaner than diesel. Biodiesel is mainly composed by a mixture of fatty acid alkyl

esters (FAAE), which can be produced from vegetable oils, wasted cooking oils, and

animal fats. Thus it is considered as renewable fuel source. Recently biodiesel has been

promoted as a way for enhancing energy independence, promoting rural development,

and reducing green house gas emissions.

2. The Glycerol’s World 11

Biodiesel can be produced through the reaction between feedstock oil with either

methanol or ethanol. Oil’s solubility in methanol is lesser than in ethanol, and rate reaction

is mass transfer limited and methanol makes higher equilibrium conversion due to higher

reactive intermediate methoxide. Most of the biodiesel is produced currently using

methanol, which is petrochemically obtained. This dependence on methanol could be

considered as non renewable basis. On this way, different efforts to produce biodiesel

from ethanol are carried out to generate a renewable process [3-5]. Also ethanol could be

a renewable alternative to produce biodiesel because it can be obtained from glycerol

which could be also obtained during the same biodiesel process [3-6].

Transesterification process can be carried out by two ways, chemically or biocatalytically

catalyzed. Chemical catalysis has other two alternatives, alkali- and acid- catalysis.

Industrial production of fuel biodiesel is performed by methanolysis using alkaline

catalysts, and high conversion levels in short reaction times are reached. However this

way has several drawbacks: free fatty acids (FFAs) and water interfere with the reaction

generating fatty acid alkaline salts (soaps). Soaps should be removed by washing water,

which also removes glycerol, methanol (MeOH), and catalyst. Also alkaline catalyst has to

be removed from the product. Raw glycerol as by-product should be treated as a waste

material making the glycerol recovery difficult, and the alkaline wastewater requires

treatment. It is also an energetically intensive process [7-8]. On the other hand, in acid

catalysis process, sulfuric and sulfonic acid are preferred because these carry out high

alkyl esters yields. But elevated reaction temperatures (>100°C) and reaction times (ca 50

h) to complete conversion are required.

Commonly, a catalyst is used to improve reaction rate and yield, and an alcohol excess is

utilized to shift the equilibrium towards the products side. Among the used catalysts are

alkalis (NaOH, KOH, sodium and potassium alcoxides, carbonates, etc), acids (sulfonic

acids, HCl, H3PO4, H2SO4, zeolites), enzymes (lypases), and whole cells.

Acid catalysis produce high alkyl esters yield. High reaction temperatures and reaction

times to obtain complete conversion are required. Basic catalysis is a quick reaction, with

high yields, which take place under moderate conditions in comparison with acid

catalysts, but chemical transesterification using an alkali-catalysis has several drawbacks

like soap formation by saponification, and difficult recovery of glycerin by emulsion


development. In contrast, biocatalysts allow synthesis of specific alkyl esters, easy

recovery of glycerol, and transesterification of glycerides with high free fatty acid content

[9], and its main disadvantages are the biocatalyst cost and lower reaction rates [10-12].

Lipase enzymes have been used in biodiesel production in free form or immobilized on

some different materials such as ceramics, kaolinites, silica, etc. [13]. In general

immobilization enhances the stability of lipase due to the ability of the support material to

retain just the right quantity of water for the enzyme to remain active. Different reactive

mixtures containing water have been analyzed. For example, immobilized Rhizopus

delemar and Rhizomucor miehei lipases efficiently catalyze alcoholysis with long-chain

fatty alcohols even in the presence of 20% water [14]. However enzymatic methods have

not been industrialized because the enzymes have high price and instability [7, 11-12].

2.3 Glycerol market and its oversupply problem

Biodiesel production could be fully sustainable if ethanol is produced from glycerol, which

is the by-product in biodiesel production. Also, enzymatic transesterification can be

carried out using ethanol with low water content or azeotropic ethanol, without affecting

considerably the biodiesel production. Genetically modified E. coli [6] and E. aerogenes

[4-5] have been reported to ferment crude glycerol or pure glycerol to ethanol. In order to

close the renewable biodiesel production in an integrated biotechnological system the

follow structure is analyzed in the next section: aqueous-ethanol as raw material,

biocatalysts use, and biological transformation of glycerol to ethanol.

Until 2003 supply of raw glycerol in the market remained relatively stable, when the

production of biodiesel started increasing in the USA [15]. Since then, the availability of

crude glycerol has been almost doubled, while its demand has remained almost

unchanged. Thus, combined effect of supply excess and limited demand of raw glycerol

led to low sale prices. Although pure glycerol is an important feedstock in many industrial

sectors, raw glycerol must be refined by large scale biodiesel producers using traditional

separation processes to remove impurities such as fatty acids, alcohol, and catalyst.

Some of these processes are filtration, chemical additions, and fractional vacuum

distillation. Generally processes are expensive and economically unfeasible for small and

medium scale plants [16].


Traditional commercial applications of glycerol are related to its use as an additive or raw

material. The industrial sectors who consume glycerol are: pharmaceutical (18%),

personal care (toothpaste and cosmetics 16%), polyether/polyols manufacture (14%),

food (11%), triacetin (10%), alkyd (8%), snuff (6%), detergents (2%), cellophane (2%),

and explosives (2%). The remaining share (11%) is used in the manufacture of lacquers,

varnishes, inks, adhesives, plastic synthetics, regenerated cellulose, and other industrial

uses [17].

Annually nearly 160000 tons of glycerol is used for technical applications and it is

expected an annual growth rate of 2.8%. Raw glycerin supply in the market remained

relatively stable until 2003, when biodiesel production started to increase in the U.S. and

the E.U. [15]. Then, availability of raw glycerin has almost doubled, and its demand has

remained largely unchanged. This excess supply and limited demand has taken to low

glycerol prices, but although refined glycerin prices have decreased in the last years; the

strongest impact has been suffered by raw glycerin, and thus its sale prices plummet

quickly [18].

2.4 Glycerol as raw material

Since 2006, the glycerol oversupply forced to biodiesel producers to receive sales prices

of 2 cents per pound or even lower prices for the raw product. But at mid-2007, reached

prices were between 6 and 10 cents per pound [19]. On the other hand, refined glycerin

prices have had a similar behavior, with prices as low as 20-30 cents per pound,

depending on the quality and purity [18-19]. In this sense the raw glycerin market will

remain weak while large amounts of this raw component being available. Therefore

glycerol is nowadays a key problem in biodiesel production, and its low sale price could

convert this by-product in a residue, then the biodiesel producers must be found

alternative uses to avoid the continue falling on the glycerol price.

Development of biorefineries based on raw glycerol to produce high-value compounds is

necessary in the biodiesel industry to overcome the economic glycerol drawback. The

simplest alternative to increase the value of raw glycerol is refining it to technical glycerin,

food or pharmaceutical grade, although to synthesize value-added components by

chemical o fermentative via are alternatives with higher potential. Chemically glycerol can


be transformed to: oxidation products on metallic catalysts as Pt, Pd, Au, using promoters

as Bi and Pb; glycols by hydrogenolysis on Ru, Cu and Pt catalysts; polyglycerols by

etherification on zeolites and mesoporous materials; and syngas by pyrolysis and

gasification.

Also, due to glycerol is abundant in nature and produced by yeasts during osmoregulation

to decrease extracellular water activity [20], its wide occurrence allows to different kinds of

microorganisms metabolizing glycerol as a sole carbon and energy source, and then this

may substitute traditional carbohydrates, such as sucrose, glucose and starch, in some

industrial fermentation processes [21-23]. Glycerol can be transform by fermentative via

to 1,3-propanediol, dihydroxyacetone, succinic acid, propionic acid, ethanol, citric acid,

pigments, polyhydroxyalcanoate, and biosurfactants. [24]. The following sections review

the main technological topics related to glycerol transformation by chemical and

biochemical via.

References [1] McCoy M. Glycerine Surplus. Chem. Eng. News. 2006, 84(6):7-8

[2] Tyson, S.K., Biodiesel Handling and Use Guidelines. 2001, U.S. Department of

Energy. NREL/TP-580-30004: CO. USA. p. 17.

[3] da Silva, G.P., Mack, M., Contiero, J., Glycerol: A promising and abundant carbon

source for industrial microbiology, Biotechnol Adv 27(1) (2009), 30-39.

[4] Yazdani, S.S., Gonzalez, R., Engineering Escherichia coli for the efficient conversion

of glycerol to ethanol and co-products, Metab Eng 10(6) (2008) 340-351.

[5] Ito, T., Nakashimada, Y., Senba, K., Matsui, T., Nishio, N., Hydrogen and ethanol

production from glycerol-containing wastes discharged after biodiesel manufacturing

process, J Biosci Bioeng 100 (2005) 260–265.

[6] Dharmadi, Y., Murarka, A., Gonzalez, R., Anaerobic fermentation of glycerol by

Escherichia coli: a new platform for metabolic engineering, Biotechnol Bioeng 94 (2006)

821–829.


[7] Fukuda, H., Kondo, A., Noda, H., Biodiesel fuel production by transesterification of oils,

J Biosci Bioeng 92(5) (2001) 405-416.

[8] Shimada, Y., Watanabe, Y., Sugihara, A., Tominaga, Y., Enzymatic alcoholysis for

biodiesel fuel production and application of the reaction to oil processing, J. Mol. Catal. B:

Enzym. 17(3-5) (2002) 133-142.

[9] Shah, S., Sharma, S., Gupta, M.N., Biodiesel preparation by lipase-catalyzed

transesterification of jatropha oil, Energ Fuel 18 (2004) 154-159.

[10] Fukuda, H., Kondo, A., Noda, H., Biodiesel fuel production by transesterification of

oils, J Biosci Bioeng 92(5) (2001) 405-416.

[11] Xu, Y., Du, W., Liu, D., Zeng, J., A novel enzymatic route for biodiesel production

from renewable oils in a solvent-free medium, Biotechnol Lett25 (2003) 1239-1241.

[12] Shimada, Y., Watanabe, Y., Samukawa, T., Sugihara, A., Noda, H., Fukuda, H.,

Tominaga, Y., Conversion of vegetable oil to biodiesel using immobilized Candida

antarctica lipase, J. Am. Oil Chem. Soc. 76 (1999) 789-793.

[13] Yagiz, F., Kazan, D., Akin, N., Biodiesel production from waste oils by using lipase

immobilized on hydrotalcite and zeolites, Chem. Eng. J. 134 (2007) 262-267.

[14] Shimada, Y., Watanabe, Y., Sugihara, A., Tominaga, Y., Enzymatic alcoholysis for

biodiesel fuel production and application of the reaction to oil processing, J. Mol. Catal. B:

Enzym. 17(3-5) (2002) 133-142.

[15] Ott, L., Bicker, M., Vogel, H., 2006. Catalytic dehydration of glycerol in sub- and

supercritical water: a new chemical process for acrolein production. Green Chem. 8, 214-

220.

[16] Posada, J.A. and C.A. Cardona, Análisis de la refinación de glicerina obtenida como

co-producto en la producción de biodiesel. Ingeniería y Universidad, 2010. 14(1)

[17] Pagliario, M. and M. Rossi., The future of Glycerol: New usages for a versatile raw

Material. 2008, Cambridge: RSC Publishing.


[18] Rahmat, N., A.Z. Abdullah, and A.R. Mohamed, Recent progress on innovative and

potential technologies for glycerol transformation into fuel additives: A critical review.

Renewable and Sustainable Energy Reviews. 14(3): p. 987-1000.

[19] Ito, T., et al., Hydrogen and Ethanol Production from Glycerol-Containing Wastes

Discharged after Biodiesel Manufacturing Process. J. Biosci. Bioeng., 2005. 100(3): p.

260-265

[20] Wang, Z., et al., Glycerol production by microbial fermentation: A review.

Biotechnology Advances, 2001. 19(3): p. 201-223.

[21] Solomon, B.O., et al., Comparison of the energetic efficiencies of hydrogen and

oxychemicals formation in Klebsiella pneumoniae and Clostridium butyricum during

anaerobic growth on glycerol. Journal of Biotechnology, 1995. 39(2): p. 107-117.

[22] Barbirato, F. and A. Bories, Relationship between the physiology of Enterobacter

agglomerans CNCM 1210 grown anaerobically on glycerol and the culture conditions.

Research in Microbiology. 148(6): p. 475-484.

[23] Menzel, K., A.P. Zeng, and W.D. Deckwer, High concentration and productivity of

1,3-propanediol from continuous fermentation of glycerol by Klebsiella pneumoniae.

Enzyme and Microbial Technology, 1997. 20(2): p. 82-86.

[24] da Silva, G.P., M. Mack, and J. Contiero, Glycerol: A promising and abundant carbon

source for industrial microbiology. Biotechnology Advances. 27(1): p. 30-39.

3. Methodology for Processes Design and Analysis

This chapter describes a methodological procedure in order to design and assess

technological schemes for the conversion of raw glycerol to added-value products. This

methodology uses a strategy based on knowledge which employs both heuristic rules and

researchers’ experience. Also it is equally applied to chemical or biochemical processes,

as well as conventional technologies or integrated process. Also, directions are given to

perform both steps: the process simulation and the process assessment.

3.1 Processes design

This thesis aims to design and assess technological schemes for the conversion of raw

glycerol obtained during the biodiesel production to added-value products. Thus, different

possibilities of glycerol conversion to added-value products should be first indentified

based on the reported literature. In this way, two main routes for glycerol transformation

are available, chemical conversion and fermentative transformation. Glycerol can be

chemically transformed by many ways such as: oxidation, hydrogenolysis, etherification,

pyrolysis, and gasification. In this sense, many catalysts such as: Pt, Pd, Au, Ru, Cu, Pt,

zeolites, and mesoporouses materials, have been widely reported for glycerol conversion.

Otherwise, many wild and metabolically engineered strains have been analyzed for the

glycerol uptake as substrate in order to produce a wide spectrum of metabolites such us:

1,3-propanediol, ethanol, poly-3-hydroxibitirate, lactic acid, propionic acid, succinic acid,

and rhamnolipids.

In order to achieve this objective it was required to: (i) classify all the information available

on glycerol transformation by chemical or fermentative routes; (ii) organize these

information based on the specific used way for either route, chemical (e.g., oxidation,

hydrogenolysis, etherification, pyrolysis, and gasification) or fermentative (e.g., production

of: 1,3-propanediol, ethanol, poly-3-hydroxibitirate, lactic acid, propionic acid, succinic


acid, and rhamnolipids); (iii) compare and analyze each transformation possibility based

on operational criteria such as: conversion, yield, and productivity; and (iv) choose the

conversion possibilities with the higher potential to be commercialized. This first stage

corresponds to both the literature review and the choosing of the most attractive

possibilities for glycerol conversion to added-value components. Moreover, it was found

that not only pure glycerol has been used as feedstock to its transformation but also crude

glycerol has been widely analyzed. Thus, in order to homogenize the feedstock used for

this study, raw glycerol obtained from the biodiesel production process was considered as

the unique feedstock. In this way, the influences of several feedstocks used for biodiesel

production were analyzed on the composition of the glycerol layer and an average

composition for the raw glycerol stream was chose. Due to this raw glycerol stream

contains low quantity of glycerol, a purification process was analyzed in order to obtain

the three most important qualities of commercially available glycerol. Under this view of

point, only one feedstock (i.e., raw glycerol) is always considered and different qualities of

glycerol (crude glycerol, technical glycerol and USP glycerol) can be used for its

transformation. Additionally, the fact of work with raw glycerol as the unique feedstock,

allows considering any designed process as an adjacent biorefinery to the biodiesel

production process.

Because of many final products and transformation routes are considered though out this

study, each alternative requires both a specific process analysis and a process design

using a strategy based on knowledge.

The synthesis of technological schemes by mean of a strategy based on knowledge

allows generating systematically alternatives which consider the specific characteristics of

each process. Thus, it is possible to design technological configurations of high

performance considering mainly techno-economic criteria. Here, the traditional

hierarchical decomposition methodology based on the onion diagram (see Figure 3.1.) for

process design is applied [1]. This sequential procedure allows designing and comparing

different alternatives for the same objective.

The process design starts analyzing the reaction step which is the fundamental stage in

this study, and then the analysis continues to the external layer of the onion diagram

adding stages such us: the separation and recycle system according to the Figure 3.1.

3. Methodology 19

This heuristic and hierarchical methodology emphasizes in both the decomposition and

analysis of different process alternatives, allowing a quick selection of technological

configurations that are often close to the best solution. Furthermore, the nature of this

approach, allows discarding many configurations easily which in general do not lead to

"good" designs. In addition, tiered design allows the use of process simulators and thus

the process diagram can be completed in an evolutionary manner. This methodology has

been applied primarily to processes of chemical or petrochemical industry.

Figure 3.1. Hierarchical decomposition according to the "onion diagram"

On the other hand, for the downstream process design the method so called breadth-first

was applied in order to analyze different alternatives for the products recovery. This

method allows both screening the best alternative for a specific purpose of the

downstream process, and evaluating of process alternatives at the next level of

hierarchical decomposition (see Figure 3.2).

Most of the alternatives for glycerol conversion to added-value compounds are analyzed

during first hierarchical decomposition levels (1 and 2) by mean of the economic potential

criteria which also involves operational variables such as conversion and yield. Thus, the

alternatives with the highest economic potential are selected to continue the synthesis of

technological schemes, while the alternatives showing unfavorable economic potential are


discarded. Thus, base structures are obtained for the chosen conversion possibilities

which are later complemented through detailed process information for the main stages of

processing.

Figure 3.2. The process design method based on the so called breadth-first

The last hierarchical levels of analysis (3-6) require more detailed information for the main

process variables which are performed by mean of sensitivity analysis and subsequent

rigorous economic evaluation. Details of the processes simulation and economic

assessment are given below.

3.2 Processes simulation Aspen Plus (Aspen Technology, Inc., USA) is the main used tool for defining, structuring,

specifying, and simulating the technological schemes for either chemical or biochemical

conversion of glycerol to added-value components.

Information required for simulating the most basic technological schemes such as:

physical and chemical properties, parameters of design, and operation of processing

units, are mainly obtained from secondary sources (e.g., articles, technical reports,

databases, patents, among others). Then, the most complex and detailed technological

schemes are obtained by mean of rigorous simulations which involve sensitivity analysis

and search of optimal operation conditions.

Because of both the petrochemical character of Aspen Plus and its modular-sequential

approach, there are not available kinetic models describing the biotechnological

processes such as fermentations or enzymatic reactions. Therefore, it is required to work

with the available interface between Aspen Plus and Excel. Additionally, the study of

3. Methodology 21

complex fermentation kinetics describing inhibition phenomena, in some cases requires

generating more complex calculation routines using other kind of softwares (e.g.,

MatLab).

Although several possibilities for glycerol conversion to added-value products have been

reported, a few publications describe kinetic models fitted good enough. Therefore, the

stoichiometric approach is here considered as a completely valid and relevant approach

for analyzing the reaction stage of different technological schemes.

On the other hand, specific compounds involved in the different processes of raw glycerol

conversion to added-value products such as: free fatty acids, alkyl esters, proteins, salts,

cell mass strains, enzymes, and other complex molecules produced by reactive-extractive

process are not available on the Aspen Plus Database. Thus, these compounds should

be created for each simulation as follows: conventional (by mean of group contribution

methods), solids (e.g., biomass), or non-conventional (e.g., enzymes).

All processes are designed and analyzed using the same calculation base which is 1000

Kg/h of raw glycerol always fed to the glycerol purification process. As the simulation

results, mass and energy balances are obtained for the technological schemes. Thus, it is

possible to obtain requirements of additional raw material, solvents, utility fluids, and

energy.

The analysis of conventional separation methods in the distillation process was carried

out with the help of the corresponding modules of the process simulators. For this, both

short-cut methods and rigorous models available in the simulation package were

employed. For simulation of the different technologies involving the operation of

distillation, the short-cut method DSTWU incorporated in the package Aspen Plus was

applied. This method uses the equations and correlations of Winn-Underwood-Gilliland in

order to provide an initial estimation of the minimum number of theoretic stages, minimum

reflux ratio, location of the feed stage, and components distribution. The rigorous

calculation of the operating conditions in the distillation columns was performed using the

module RadFrac based on the equilibrium method that employs the MESH equations

(Mass balance equations, phase Equilibrium equations, Summation of the compositions,

and Heat balance equations) using the inside-out algorithm. Residue curve maps were


used for the conceptual design of the distillation schemes applying the principles of

topological thermodynamics (analysis of the statics) [2]. Sensitivity analyses were

performed in order to study the effect of the main operating variables (reflux ratio,

temperature of the feed stream, ratio between the distillate and the feed, etc.) on the

biodiesel purity and the energy consumption of this operation. The final result is the

determination of operating conditions that allow developing energetically efficient

processes. The objective of this procedure was to generate the mass and energy

balances from which the requirements for raw materials, consumables, service fluids and

energy needs are calculated.

On the other hand, because of the significant differences involved in the reaction of

glycerol to different products, the reaction conditions are specific for each technological

scheme. Besides, the downstream process is designed based on the products distribution

obtained after the reactive stage. Thus, detailed information about reaction stage and

downstream process is given according to each case of study.

Estimation of the energy consumption is performed based on the results of the mass and

energy balances generated during the simulation process. Thus, the thermal energy

required in the heat exchangers and re-boilers was taken into account, as well as the

electric energy needed by pumps, compressors, mills, and other equipments. The energy

demand was calculated from the mass and energy balances generated by the simulator.

The balances included the energy consumption of reboilers and condensers used in

distillation columns, and the energy consumption of the reactors.

3.3 Processes assessment The capital and operating costs were calculated using the software Aspen Icarus Process

Evaluator (Aspen Technologies, Inc., USA). This software estimates the capital costs of

process units as well as the operating costs, among other valuable data, utilizing the

design information provided by Aspen Plus and the data introduced by the user for

specific conditions such as project location among others. Also, analyses are based on

the strategy designed by Cardona et al [3-7] for process assessment.

3. Methodology 23

This analysis was estimated in US dollars for a 10-year period at an annual interest rate

of 16 %, considering the straight line depreciation method and a 33% income tax [8]. The

cost for raw glycerol, crude glycerol, and refined glycerol as well as the labor cost for

operatives and supervisors, and the prices for electricity, water and low pressure vapor

are showed in the Table 3.1. Additionally, the commercial price for other required

compounds such us raw materials and solvents are listed in the Table 3.1.

Table 3.1. Used costs and prices for the economic assessment

Costs Value Units

Operatives 2.14 UDS$/h

Supervisors 4.29 UDS$/h

Electricity 0.03044 UDS$/kwh

Water 1.252 UDS$/m3

Low pressure vapor 8.18 UDS$/Ton

Raw glycerol 132.45 UDS$/Ton

Crude glycerol (85 wt %) 540.84 UDS$/Ton

Refined glycerol (98 wt %) 706.41 UDS$/Ton

Succinic acid 2492.2 UDS$/Ton

Lactic acid 1552.2 UDS$/Ton

Acetic acid 591.8 UDS$/Ton

Dichloromethane 850 UDS$/Ton

Trioctylamine 2550 UDS$/Ton

Methanol 290 UDS$/Ton

1-Octanol 1835 UDS$/Ton

Iso-butylaldehyde XX UDS$/Ton

1,3-Propanediol 1766 UDS$/Ton

Propionic acid 1220 UDS$/Ton

DES 3050 UDS$/Ton

Glucose 480 UDS$/Ton

PHB 3500 UDS$/Ton

Propionic acid 1800 UDS$/Ton


References

[1] Smith R.M. The nature of chemical process design and integration. Chemical Process:

Design and Integration. John Wiley & Sons: Hoboken, NJ, USA, (2005), pp 1-15.

[2] Pisarenko Y.A., Serafimov L.A., Cardona C.A., Efremov D.L., Shuwalov A.S., Reactive

distillation design: analysis of the process statics. Reviews in Chemical Engineering 17

(4), 2001.

[3] Cardona, C.A., Sánchez, O. J., Energy consumption analysis of integrated flowsheets

for production of fuel ethanol from lignocellulosic biomass. Energy. 2006. 31:2447-2459.

[4] Cardona, C.A., Sánchez, Ó.J., Fuel ethanol production: Process design trends and

integration opportunities. Bioresource Technology. 2007. 98 2415-2457.

[5] Quintero, J.A., Montoya, M.I., Sánchez, O.J., Giraldo, O.H., Cardona, C.A., Fuel

ethanol production from sugarcane and corn: Comparative analysis for a Colombian case.

Energy. 2008. 33:385-399.

[6] Cardona, C.A., Gutiérrez, L.F., Sánchez, O. J., In Energy Efficiency Research

Advances (D. M. Bergmann, ed.). 2008. Pp:173-212. Nova Science Publishers,

Hauppauge, NY, USA.

[7] Gutiérrez, L.F., Sánchez, Ó.J., Cardona, C.A., Process integration possibilities for

biodiesel production from palm oil using ethanol obtained from lignocellulosic residues of

oil palm industry. Bioresource Technology. 2009. 100:1227-1237.

[8] Posada J.A., Cardona C.A., Rincón L.E., Sustainable biodiesel production from palm

using in situ produced glycerol and biomass for raw bioethanol. In 32nd

symposium on

biotechnology for fuels and chemicals. Clearwater Beach, Florida. April 19-22. 2010.

4. Separation and Purification of Glycerol

This chapter presents commercial, technical, and technological aspects related to the

glycerol purification process such as the most important commercially qualities of glycerol,

the influence of the feedstock used for biodiesel production on the glycerol layer, and the

conventional and non-conventional purification processes. The most important qualities of

commercial glycerol are: crude glycerol (80-88 wt %), technical glycerol (98 wt %), and

refined glycerol (USP or FCC grades, 99.7 wt %). Thus, a flowsheet able to purify raw

glycerol up to these three qualities was designed, simulated, and economic assessed.

Simulation results showed that is possible to reach the quality requirements while

economic results showed that is a profitable process. Also, recovering of anhydrous

methanol at 99 wt % could represent an additional incoming for the purification process

which could reduce the purification costs among 19 to 26 %. Simulation process is carried

out using Aspen Plus software, while the economic evaluation is performed by Aspen

Icarus Process Evaluator package.

4.1 Commercial qualities of glycerol Most of the marketed glycerol is manufactured to satisfy the strict requirements of the

United States Pharmacopeia (USP) and the Food Chemicals Codex (FCC) (The Soap

and Detergent Association). However, technical grades of glycerol which have not been

certified as USP or FCC are also available in the market. The three main qualities of

glycerol commercially available depend on their purity, these are: raw glycerol, technical

glycerol, and refined glycerol (USP or FCC grade). Raw glycerol usually has between 40

and 88 wt % of glycerol, and contains high amount of methanol, soaps, and salts. This

glycerol is commonly obtained as by-product on biodiesel production. Technical glycerol

is a high purity product where most of its pollutants have been totally removed. This

glycerol is free of methanol, soaps, salts, and other components. Refined glycerol is a

pharmaceutical quality product which can be used in foods, personal care, cosmetics,


pharmaceutical products, and other special applications. Also, these products must

complete the specifications of Pharmacopeia of the USA (USP 30) and the Food and

Drug Administration (FDA) of the USA. Table 4.1 shows the main quality specifications

and the thresholds for the pollutants present in this glycerol [1].

Table 4.1. Quality specifications for the main qualities of glycerol

Properties

Raw Glycerol

Technical Glycerol

Refined Glycerol (USP)

Glycerol Content 40-88% 98.0% Min 99.70%

Ash 2.0% Max N/A N/A

Moisture N/A 2.0% Max 0.3% Máx.

Chlorides N/A 10 ppm Max 10 ppm Máx.

Color N/A 40 Max (Pt - Co) 10 Max. (APHA)

Specific Gravity N/A 1.262 (@25°C) 1.2612 Min

Sulfate N/A N/A 20 ppm Máx

Analysis N/A N/A 99.0 - 101.0% (dry base)

Heavy Metals N/A 5 ppm Máx. 5 ppm Máx.

Chlorates Components

N/A 30 ppm Máx. 30 ppm Máx.

Ignition Residues N/A N/A 100 ppm Máx.

Fatty acids and Esters

N/A 1.00 Máx 1000 Máx

Water 12.0% Max 5.0% Máx 0.5% Máx

pH (solution 10%) 4.0 - 9.0 4.0 - 9.1 N/A

Organic Residues 2.0% Máx 2.0% Máx N/A

4. Glycerol Purification 27

4.2 Effect of the feedstock for biodiesel production on glycerol composition

Raw glycerol has a very low value in the market because of its impurities. Also,

composition of glycerol highly depends on both the family of used raw material and the

process conditions for biodiesel production. This fact occurs because the chemical

compositions of the feedstocks used for biodiesel production could change significantly.

Fats and oils usually contain more than ten types of fatty acids, which have between 12

and 22 carbons. But, often the higher proportion of fatty acids has between 16 and 18

carbons. Although these fatty acids are saturated, monounsaturated or polyunsaturated

[2], different degrees of saturation affect the properties of the biodiesel fuel. Thus, a

"perfect" biodiesel should be only obtained from monounsaturated fatty acids.

The composition profile of fatty acids was presented by He and Thompson [3] for six

vegetable oils (i.e., IdaGold mustard, PacGold mustard, rapeseed, canola, crambe, and

soybean) and for waste vegetable oil (WVO) used as feedstocks on biodiesel production

as shown in Table 4.2. Additionally, based on the reported information by He and

Thompson [3], the composition of glycerol layer was calculated for each used feedstock

for biodiesel production. The results are shown in Table 4.3.

Table 4.2. Fatty acid profile of vegetable and used oils [3]

Composition (wt %)

Fatty acids IdaGold PacGold Rape Canola Soybean Crambe WVO

Palmitic (16:0) 2,8 3,1 2,9 4,5 10,7 2 18,7

Estearic (18:0) 1 1,6 1 1,8 4,3 0,9 6,3

Oleic (18:1) 24,8 23,9 13,7 60,7 24,9 17,9 40,5

Linoleic (18:2) 10,3 21,6 11,8 19,1 51,6 8,1 28

Linolenic (18:3) 9,4 9,9 7,5 9,5 7,3 4,5 1,5

Eicosic (20:1) 10,7 12,1 8,7 1,8 0,2 3,7 ---

Erucic (22:1) 34,7 22,1 48,5 0,9 --- 54,1 ---

MW Average (Kg/Kmol) 946,3 924,6 968,5 882,1 872,8 978,5 867,2


Table 4.3. Composition of the glycerol layer obtained by decantation during the biodiesel

production from different feedstocks

Oil

Component IdaGold PacGold Colza Canola Soja Crambe WVO

Methanol (wt %) 32,59 32,68 28,20 25,07 26,06 23,17 11,72

Glycerol (wt %) 60,05 61,39 59,94 60,38 61,67 65,01 46,41

NaOCH3 2,62 (wt %) 2,82 2,27 2,24 2,56 2,69 1,99

Proteins (wt %) 0,13 0,18 0,06 0,05 0,05 0,46 0,14

Fats (wt %): 1,94 1,08 8,88 11,68 7,17 8,42 36,41

Palmitic (16:0) 0,054 0,030 0,249 0,327 0,201 0,236 1,020

Estearic (18:0) 0,019 0,011 0,089 0,117 0,072 0,084 0,364

Oleic (18:1) 0,480 0,269 2,203 2,896 1,779 2,087 9,030

Linoleic (18:2) 0,200 0,112 0,915 1,203 0,739 0,867 3,750

Linolenic (18:3) 0,182 0,102 0,835 1,098 0,674 0,791 3,423

Eicosic (20:1) 0,207 0,116 0,951 1,250 0,767 0,901 3,896

Erucic (22:1) 0,672 0,376 3,082 4,052 2,489 2,921 12,635

Ash (wt %) 2,67 1,85 0,64 0,58 2,48 0,26 3,33

4.3 Conventional purification process At laboratory scale the purification of the system containing biodiesel, glycerol, soaps, and

salts (mainly sodium methoxide, NaOCH3), is preformed using separation funnels, which

allow to soaps being remained in the crude glycerol layer. Layer containing esters must

be heated up to 85 °C in order to recover the unreacted methanol. While industrially raw

glycerol is refined through a filtration process, followed by mixing with chemical additives

which allow the precipitation of salts and finally different qualities of commercial glycerol

are obtained by a vacuum fractional distillation process.


Distillation is the most commonly used method for glycerol purification. This technology

produces high purity glycerol at high yields. However, the glycerol distillation is an energy

intensive process because of its high heat capacity, requiring a high supply of energy for

vaporization [4]. Ion exchange has also been used to purify raw glycerol [5], but this

technique is not economically viable from an industrial view of point due to the high

content of salts. Also, when contents of sales are above 5 wt % which is tipically found in

the glycerol stream obtained from the biodiesel industry, the chemical regeneration cost of

these resins becomes very high. Figure 4.1 shows the flow diagram for the two above

described conventional techniques for glycerol purification.

P-7

Filtration

Concentration by evaporation

Evaportation and refining of crude glycerol

Purification by ionic exchange

Purification by vaccum distillation

Figure 4.1. Flowsheet of conventional schemes for glycerol purification.

4.4 Alternative purification processes A commercially available technology for raw glycerol purification obtained during the

biodiesel production was jointly developed by Rohm and Haas, a provider of functional

polymers by ion exchange technologies and catalysts, and by Novasep Process, a

supplier of purification solutions which includes chromatography, ion exchange,

membranes, crystallization, and evaporation. The process is the so called Ambersep

BD50 [6]. In principle, this process uses a chromatographic separator in order to remove

large amount of salts and free fatty acids. Refined stream is then processed in an

evaporator / crystallizer unit, which removes the salts in a crystalline form. This fact


avoids the effluents production in the glycerol purification plant. Thus, a glycerol stream at

a purity of 99.5 wt % is obtained. But if a high quality glycerol is required, (e.g., 5 to 10

parts per million of salt content) it is possible to use a ion exchange demineralization unit.

This process has lower energy requirements compared to the traditional distillation

process. The block diagram for the Ambersep BD50 process is shown in Figure 4.2 which

illustrates the different steps for the raw glycerol purification process.

Figure 4.2. Flowsheet of the Ambersep BD50 process

Pre-Heating

Heating up to 90ºC

Filtration

Degasification

CHR

Heating up to 90ºC

Filtration

Degasification

Cooling at 40ºC (Optional)

IEX (Optional)

Concentration

Crystallization NaCl

NaCl Secondary Glycerol

Refined

Effluent

Crude

Refined Glycerol

Light Water


4.5 Simulation of the glycerol purification processes Based on the calculated compositions for the glycerol layer obtained from different

feedstocks (see Table 4.3.), the profile compositions obtained from IdaGold mustard

represents the average values among the first use oils analyzed. Thus, this stream was

chosen to design the purification process of the raw glycerol.

Figure 4.3.a. shows the simplified flowsheeet for raw glycerol purification to 88 wt %

(crude glycerol) and to 98 wt % (technique glycerol). In order to obtain glycerol at 99.7 wt

% (glycerol USP grade), it is required a further refining process throughout a ion

exchange resin which removes the triglycerides still contained in the mixture, as shown in

Figure 4.3.b.

a

1 2 3 4 5 6

Raw Glycerol

Methanol

Solids

Water

OrganicPhase

Aqueous Glycerol

Water waste 1

Water waste 2

Glycerol

b

RII-1 Glicerina USP

Adsorbato

Figure 4.3. Simplified flowsheet for raw glycerol purification. a) purification at 88 and 98

wt %. b) Purification at 99.7 wt %. 1. First evaporation column, 2. Neutralization tank, 3.

Centrifuge, 4. Decantation tank, 5. Second evaporation column, 6. Distillation column. 7:

Ionic exchange resine.

The raw glycerol stream is initially evaporated, where 90 % of methanol at 99 wt % is

recovered. Also, since glycerol is the unique impurity present in the recovered stream, this

steam of anhydrous methanol is appropriate to be reused in the transesterification

process. Bottom stream obtained from Evaporator I is neutralized using an acid solution.

Then both salts produced during the neutralization process and remaining ashes and

proteins are retired by centrifugation. The clarified product obtained from the centrifuge is

washed with water using a weight ratio of 2.4 (water/glycerol stream). Thus, 50% of the

triglycerides remaining in the mixture are withdrawn with a glycerol lost of 1.8 %. The

resulting aqueous glycerol stream and free of salts, solids, and protein but with a low

content of both methanol and triglycerides, is subjected again to an evaporation process

which removes more than 90% of water and most of the remaining methanol, with a


glycerol lost of 0.2%.Thus, the glycerol purity reached is 80 wt %. Then, the glycerol

stream is purified through a distillation column to reach the required purity, either 88 wt %

or 98 wt %. Although in all cases the used flowsheet is the same, the operational

conditions change depending on the required purity.

In order to obtain glycerol at USP grade, the process conditions adjusted for glycerol at 98

wt % are in general preserved, but both the reflux ratio and the ratio of distillate/feed are

increased for the distillation tower. Also, a final refinement stage through an ion exchange

resin is required to remove 95 % of the triglycerides contained still in the mixture. Table

4.4 summarizes the simulation results obtained for the purification processes of raw

glycerol. Also, it can be observed that the obtained products meet the quality

requirements shown in Table 4.1.

Table 4.4. Simulation results for raw glycerol purification process.

Variable

Streams

Raw Glycerol

Methanol

Glycerol at 88%

Glycerol at 98%

Glycerol at 99,7%

Temperature (ºC) 25 144,2 104,7 189,2 204

Mass flow free of ash (kg/hr)

973,3

301,981

665,25

596,595

586,179

Mass fraction: 0,014

Triglycerides 0,02 0 0,015 0,016 0,001

Methanol 0,335 0,99

Water 0 0 0,105 0,004 0,002

Glycerol 0,617 0,01 0,88 0,98 0,997

NaOCH 0,027 3 0 4,3 ppm 4,3 ppm 1,62 ppm

Protein 0,001 0 0,2 ppm 0,2 ppm 75 ppb

Mass flow of ash (kg/hr) 27,6 0 0,003 0,003 0,00015


4.6 Economical assessment for glycerol purification processes

Because of the fed glycerol stream contains 32.6 wt % of methanol it is required to

consider two different scenarios. The first one considers that the withdrawn methanol is

not recovered while the second scenario considers that the withdrawn methanol from the

raw glycerol stream is recycled and reused as feedstock during the transesterification

process since this stream is composed of 99 wt % methanol and 1 % glycerol. Thus,

under the light of the second scenario, methanol is considered as a by-product stream

which has an economical value. Economic assessment results for raw glycerol purification

to 88, 98, and 99.7 wt % are shown in Table 4.5. The purification costs (PC) for each

purification process are in the first column discriminated by raw materials, utilities,

operating labor, maintenance and operating charges, plant overhead, general and

administrative costs, capital depreciation, and co-products credit. The second column

contains the share of each item by each purification process.

Table 4.5. Purification costs (PC) of raw glycerol (US$/L)

Item

(US$/L)

CP Glycerol

88 wt %

% CP (88%)

CP Glycerol

98 wt %

% CP

(98%)

CP Glycerol

99.7 wt %

% CP

(99.7%)

Raw materials 0.05539 24.78 0.05539 23.55 0.05539 23.11

Utilities 0.03741 16.73 0.07290 31.00 0.13544 56.51

Operating labor 0.01889 8.45 0.01889 8.03 0.02173 9.07

Maintenance 0.00721 3.22 0.00793 3.37 0.00979 4.08

Operating charges 0.00472 2.11 0.00520 2.21 0.00543 2.27

General costs 0.01305 5.84 0.01436 6.11 0.01576 6.57

Administrative costs 0.01093 4.89 0.01179 5.01 0.01257 5.24

Capital depreciation 0.07595 33.97 0.07595 32.30 0.08983 37.48

Co-products credit 0.05900 -26.39 0.06574 -25.05 0.06691 -19.34

CT without sale of metanol 0.22356 0.26241 0.34593

CT with sale of metanol 0.16456 0.19666 0.27902


Due to the low commercial prices of raw glycerol obtained from the biodiesel production,

raw materials represent less than 25% of the total purification cost. The capital cost

accounts the most share of the purification cost, being it about 35%. Moreover, an

increase in the glycerol purity represents an increase in the utilities cost, reaching up to

56.5% of total purification cost when glycerol at 99.7% is obtained. On the other hand, the

recovery of anhydrous methanol at 99 wt % represents a significant reduction in the total

purification cost, with a decreasing between 19 – 26 % of the total costs.

The purification cost of raw glycerol obtained from biodiesel production was reported by

Johnson and Taconi [7] as 0.15 USD$/lb or 0.26 USD$/L. This value is close to the total

purification cost obtained for glycerol at 98% at the scenario that no considers the sale of

anhydrous methanol as co-product. This scenario is the most standard analyzed since it

is a technical quality of glycerol with no by-products production.

Commercial sale prices for different qualities of glycerol are as follows: 0.28 USD$/L for

glycerol at 88 wt %, 1.39 USD$/L for vegetable glycerol at 98 wt %, 1.11 USD$/L for

tallow glycerol at 98 wt % and 3.48 USD$/L for glycerol USP grade or at 99.7 wt %. For

the assessed production scale, the purification and refining of raw glycerol is profitable

since the purification costs are lower than their selling prices.

4.7 Conclusions Commercially three qualities of glycerol were identified as the most important ones. Crude

glycerol with a purity ranging from 80-88 wt %, technical glycerol mainly found at 97 wt %,

and refined glycerol (USP or FCC grades) at 99.7 wt %. These three types of glycerol

differ significantly in the content of water, fatty acid residues, esters, and other organic

wastes. Also, some differences were found for the use of diverse feedstocks for biodiesel

production on the composition of the glycerol layer. Although, most of the first use oils

lead to not big differences in the glycerol layer, a completely different behavior was

observed for the glycerol obtained from WVO represented by low concentration of

glycerol and methanol with a high content of fats. On the other hand, based on the

traditional purification of glycerol, a flowsheet able to purify raw glycerol up to the three

commercial qualities above described was designed, simulated and economically

assessed. Results showed that not only quality requirements were successfully obtained


but also for the analyzed purification scale all the processes were profitable. Thus, a

homogenized raw material and purification process was obtained in order to continue the

analysis of different possibilities of glycerol transformation to added-value products.

References

[1] Posada, J.A. and C.A. Cardona, Análisis de la refinación de glicerina obtenida como

co-producto en la producción de biodiesel. Ingeniería y Universidad, 2010. 14(1)

[2] Belén-Camacho, D.R, Sánchez, E.D., García, D., Moreno-Álvarez, M.J., Linares O.

Características fisicoquímicas y composición en ácidos grasos del aceite extraído de

semillas de tomate de árbol (Cyphomandra betacea Sendt) variedades roja y amarilla.

Grasas y Aceites. 2004, 55(4):428-433

[3] Thompson, J.C., He, B.B., Characterization of crude glycerol from biodiesel production

for multiple feedstocks. Appl. Eng. Agric. 2006, 22(2):261-265.

[4] Posada, J.A., Cardona, C.A., Rincón L.E., Sustainable biodiesel production from palm

using in situ produced glycerol and biomass for raw bioethanol. En: Society for Industrial

Microbiology. 32nd

[5] Berriosa, M., Skelton, R.L., Comparison of purification methods for biodiesel. Chem.

Eng. J. 2008, 144:459-465.

symposium on biotechnology for fuels and chemicals. Clearwater

Beach, Florida. April 19-22. 2010.

[6] AMBERSEP™ BD50 Technology. <www.amberlyst.com/glycerol.htm> [Consulted: 29-

09-2009].

[7] Johnson, D.T., Taconi, K.A., The Glycerin Glut: Options for the Value-Added

Conversion of Crude Glycerol Resulting from Biodiesel Production. Environ. Prog. 2007,

26(4):338-348.

5. Chemical Conversion of Glycerol

This chapter presents different ways of glycerol transformation to added-value products.

Different reactions are described, such as: (i) oxidation on metallic catalysts like Pt, Pd,

Au, and on promoters as Bi and Pb; (ii) hydrogenolysis to glycols on Ru, Cu and Pt

catalysts; (iii) etherification to polyglycerols on zeolites and mesoporous materials; (iv)

pyrolysis and gasification, where the objective is to produce syn-gas.

Glycerol is a potentially important feedstock for biorefineries, available as a byproduct in

the biodiesel production by transesterification of vegetable oils or animal fats. Also, due to

its high functionality, there are many transformation ways to produce added-value

compounds using glycerol as sole feedstock for its conversion. On the other hand, new

uses for glycerol need to be found since the biodiesel production cost vary inversely with

the glycerol cost.

The high differences between the price of raw glycerol and refined glycerol, added to its

chemical versatility have carried out an intense research for developing alternative uses

and practical technologies to utilize the raw glycerol. In this sense chemical possibilities of

glycerol transformation are reviewed as follows. Thus, several transformation possibilities

to added value products have been found by chemical or biochemical ways.

5.1 Oxidation Glycerol oxidation on metallic catalysts is carried out by mean of oxidative

dehydrogenation mechanism on the metal surface [1]. First step is the alcohol

dehydrogenation, followed by the oxidation of intermediate formed [2]. Due to the

potential complexity of products distribution (Figure 5.1) the selectivity control on the

process oxidation is key [3].


The main derived oxygenated products from glycerol (GLY), are: glyceric acid (GLYAC),

dihydroxyacetone (DHA), hydroxypyruvic acid (HYPAC), tartaric acid (TARAC), mesoxalic

acid (MESOXAC), oxalic acid (OXALAC), besides some intermediates as glyceraldehyde

(GLYAL), glycolic acid (GLYCAC), and glyoxylic acid (GLYOXAC) as is shown in the

Figure. 5.1.

Figure 5.1. Possible products for glycerol oxidation

The most studied metallic catalysts are palladium (Pd), platinum (Pt), and gold (Au),

although the main disadvantage of Pd and Pt are their deactivation with the reaction time

increment [2]. To improve the activity, selectivity, and stability of the reactive system,

promoters are used on Pt and Au for redox reactions; there are particularly heavy metals

from groups IV (lead, Pb) and V (bismuth, Bi) [4].This fact allows preventing the products

over-oxidation on the metal surface, avoiding the products degradation until total oxidation

to carbon dioxide, also promoters favors the secondary alcohols oxidation. Primary

alcohols are oxidized to carboxylic acids (GLYAC, TARAC and HYPAC via DHA) in

absence of promoters or under basic pH, and secondary alcohols are selectively oxidized

on Pt-Bi metallic catalyst at acid pH (DHA, HYPAC via GLYAC and MESAC) [1-2].

Gallezot et al. [1, 5-8], Hutchings et al [9-11], Prati et al [12-18], Claus et al [2, 19] and

Davis [20-22] have studied the selective glycerol oxidation on mono - or bi - metallic

catalysts of Pd, Pt, and Au, using oxygen as oxidizer agent. Gallezot et al showed that

5. Chemical Conversion 39

GLYAC and TARAC are obtained under basic pH, while HYPAC is obtained under not

very acid pH via DHA and, DHA and HYPAC are obtained under acid pH via GLYAC and

MESAC [4, 6, 8, 23]. Total glycerol conversion is achieved for Pd and Pt catalysts with

selectivities of 70% and 35% to GLYAC and HYPAC respectively. Also, for Pt-Bi catalyst

selectivities of 83%, 74%, 37%, and 39% to TARAC, HYPAC, DHA, and MESOXAC are

obtained respectively, with conversions upper to 75%, except for MESOXAC which was

53%. On the other hand, proofs carried out with activated coal (AC) as support showed

that 5% Pd/CA catalyst has higher potential redox than 5% Pt/CA [4], and for Au catalyst

was found that activity and selectivity increase when the particle diameter diminishes.

Hutchings et al [3, 9-11] and Prati et al [12-18, 24] studied the glycerol reaction on Au

catalysts. Au supported on carbon (Au/C) is extremely selective to GLYAC (>82%), with

conversion higher to 60% [3]. Also, in systems at basic pH the selectivity to GLYAC is

increases with both, pH and oxygen pressure. Bi metallic catalysts of Pd-Au take to total

GLY conversion with high GLYAC selectivity (>45%), which was increased when bi-

metallic catalysts were immobilized on graphite.

GLY oxidation by Au catalysts supported on graphite, activated coal, and carbon

nanoparticles was studied by Claus et al [2, 19] , who found that the last one support is

the most chemically active, also confirmed the dependence among selectivity to GLYAC

and particle size. Other mono- and bi-metallic nanoparticles of Au-Pd were evaluated by

Davis et al [20-22] for GLY oxidation in liquid phase. The highest turnover frequency

(TOF) was exhibits for the Au mono-metallic catalyst and the highest selectivity to GLYAC

was reached by Pd. Also, activity and selectivity for bimetallic catalyst Au-Pd was

dependent of the Au quantity.

5.2 Reduction Glycerol reduction produces mainly 1,2- propileneglycol (12-PG), 1,3-propileneglycol (13-

PG), ethyleneglycol (ETGLY), and other by-products such as lactic acid (LACAC), acetol

(ACET), acroleine (ACRO), besides degradation products such as propanol (PROPOH) ,

methanol (METOH), methane (MET), and carbon dioxide (CO2) [25]. Among glycerol

reduction products the most important is propyleneglycol because of its high functionality

which can be used in unsaturated polyester resins, functional fluids (antifreezes and heat


transfer), pharmaceuticals, foods, cosmetics, liquid detergents, tobacco humectants,

flavors and fragrances, personal care, paints, and animal feed.

Several technological schemes for propyleneglycol production from glycerol have been

patented [26-29] in which are used different catalysts such as copper, zinc, ruthenium,

cobalt, magnesium, molybdenum, nickel, palladium and platinum; under a widely

operation conditions for pressure (2000 – 5000 psi) and temperature (200 – 350 °C). On

the other hand Shanks and Lahr [30, 31] studied the interactions among reactants and

catalyst for dehydrogenation/hydrogenation process with Ru supported on activated

carbon (5% wt Ru/CA), thus pH effects, competitive adsorption, and products degradation

(ethylene glycol and propylene glycol) were analyzed under high pressure, meddle

temperature, and high glycerol concentration (1450 psi, 205 °C and 10 wt % of glycerol).

Due to the high catalytic activity of Ru was found that ethyleneglycol and

propyleneglycerol degradation rate is independent of the initial glycol concentration,

although propylene glycol is less competitive than ethylene glycol to active sites. Also,

while selectivity to propyleneglycol is independent of pH, selectivity to ethyleneglycol

increases at low basic conditions.

Lahr and Shanks [30] purposed a model for glycerol reduction, in which glycerol is

adsorbed and dehydrogenated reversibly on the metallic catalyst where glyceraldehyde is

formed, which is then desorbed and could react through four different way in a basic

media: (i) retro-aldol mechanism to produce glycol aldehyde as precursor of ethylene

glycol; (ii) oxidation and subsequent descarboxylation to produce also glycol aldehyde;

(iii) dehydration to 2- hydroxypropionaldehyde as precursor of propylene glycol; or (iv)

degradation to unwanted side products which is also a possible way to produce the

glycols’ precursors. Finally, the respective precursors are hydrogenated to glycols. A new

mechanism for glycerol reduction under moderate operation conditions was proposed by

Dasari et al [25] where hydroxyacetone is formed by dehydrogenation of glycerol, which

after react with hydrogen to produce propylene glycol and water.

The highest selectivities to propylene glycol have been reported for Cu-based catalysts

which exhibit low selectivities to ethylene glycol and other degradation by-products [25].

While Ru- and Pd- based catalysts have low selectivities to propylene glycol (< 50%

generally) because of competitive hydrogenolysis where C-C and C-O bonds are taken to


an excessive degradation to produce lower alcohols and gases [25, 32]. In general terms

glycerol conversion is significantly increased by temperature, while yield has a maximum

near to 200°C due to the degradation products which occurs at high temperature. Also,

propyleneglycol selectivity can be improved increasing the water contend in the glycerol

mixture which reduce the glycerol conversion, but however the net yield increase.

5.3 Etherification Glycerol etherification takes to polyglycerols which are oxygenated compounds used as

surfactants, lubricants, cosmetics, and food preservatives. Polyglycerols have low

polymerization level, and these can be obtained in lineal, cyclic, or branched chains, but

researching effort s are focused on selective production of di- and/or tri-glycerols.

Selectivity of glycerol etherification is similar to pseudo-polymerization where generally a

mixture of lineal and cyclic polyglycerols is obtained, especially in presence of

homogeneous catalysts such as sodium, potassium or carbonate hydroxide [32-34].

Etherification selectivity in the first reaction step on acid catalysts is not really controlled

and a mixture of di- to hexa- glycerols (lineal or cyclic), esters of polyglycerol, and

acroleine as by-products is obtained. Although selectivity in the first step could be slightly

improved modifying the pseudo-pore size in the mesoporous materials [35]. On the other

hand, glycerol conversion was improved by Na2CO3, although low selectivities to di- and

tri-glycerols were as obtained. Then alkaline exchange zeolites were studied and

selectivity was increased [32]. Incorporation on mesoporous catalytic structure of

elements such as Al, Mg, and La, modifying only the activity, and selectivity is hold almost

constant. Clacens et al [35] found that impregnation method takes to materials most

stable and selective than incorporation method. Among the impregnated materials La is

the most active but its selectivity was the worst; a positive behavior was found to Mg

which is highly selective

5.4 Pyrolysis and gasification Pyrolysis process produces liquid fuels at temperatures between 400 – 600 °C, and gas

products at temperatures upper to 750 °C. Although gasification process is similar to


pyrolysis the main difference is that gasification is carried out in presence of oxygen like:

air, or pure oxygen, or vapor.

Reactions catalyzed by protons or hydroxyl ions can be performed under almost- or

super-critical water conditions (P> 22.1 MPa and T> 647 K) because of water is not only a

solvent it is also a catalyst due to the self-dissociation which takes to formation of

hydroxyl ions and protons. Under these conditions two competitive ways have been

identified. The first one consists in a series of ionic reactions which occur at high pressure

and/or low temperature. The second is a degradation reaction of free radicals, which

occurs at low pressure and/or high temperature. On the other hand, temperature

increases the reaction rate until critical temperature is reached, then reaction rate

decreases drastically related to subcritical conditions.

The main products of glycerol degradation are: methanol, acetaldehyde, propionaldehyde,

acroleine, allylic alcohol, ethanol, formaldehyde, carbon monoxide, carbon dioxide, and

hydrogen. Acetaldehyde and formaldehyde formation increase with the pressure which

indicates that these compounds are mainly formed by the ionic reaction, while methanol

and allylic alcohol formation decrease with the pressure which indicates that these

compounds are formed by the free radicals way [36]. Formation of gasses products

happens to high temperature; also gases formation decrease with the pressure, this

indicates a production by a reaction mechanism of free radicals. Gases formation occurs

at high temperature; also gases formation decrease with the pressure which indicates that

these are produced by a free radicals mechanism.

Syngas is the main product of pyrolysis and gasification processes. Syngas is a mixture of

hydrogen (H2) and carbon monoxide (CO). A wide range of conversions and selectivities

have been reported depending on the operational conditions such as temperature,

pressure, and glycerol concentration [36-41], also the pollutants presence for raw

glycerol, such as methanol and KOH [41]. Low glycerol concentration and high

temperature takes to high CO2 concentrations in the gas product and the most products

remain in the liquid phase [42]. On the other hand when temperature is increased the H2

and CO2 production are improved, which takes to a low CO concentration in the gas

product.


Nitrogen (N2) is used like carrier gas for glycerol pyrolysis. High amounts of N2 takes to

high liquid phase yield and gas production diminished. This process has a yield of 93% to

syngas (H2 + CO) at 800 °C like showed Valliyappan [41]. On the other hand gasification

is carried out with vapor without any carried. Total glycerol conversion was reported for a

initial mixture of 50 wt % of glycerol by Valliyappan from pure and raw glycerol [41].

References

[1] Gallezot, P., Selective oxidation with air on metal catalysts. Catalysis Today, 1997.

37(4): p. 405-418.

[2] Demirel-Gülen, S., M. Lucas, and P. Claus, Liquid phase oxidation of glycerol over

carbon supported gold catalysts. Catalysis Today, 2005. 102-103: p. 166-172.

[3] Hutchings, G.J., Catalysis by gold. Catalysis Today, 2005. 100(1-2): p. 55-61.

[4] Garcia, R., M. Besson, and P. Gallezot, Chemoselective catalytic oxidation of glycerol

with air on platinum metals. Applied Catalysis A: General, 1995. 127(1-2): p. 165-176.

[5] Fordham, P., et al., Selective catalytic oxidation with air of glycerol and oxygenated

derivatives on platinum metals, in Studies in Surface Science and Catalysis. 1996,

Elsevier. p. 161-170.

[6] Fordham, P., et al., Selective oxidation with air of glyceric to hydroxypyruvic acid and

tartronic to mesoxalic acid on PtBi/C catalysts, in Studies in Surface Science and

Catalysis. 1997, Elsevier. p. 429-436.

[7] Fordham, P., M. Besson, and P. Gallezot, Selective oxidation with air of glyceric to

hydroxypyruvic acid and tartronic to mesoxalic acid on PtBi/C catalysts. Stud. Surf. Sci.

Catal., 1997. 108: p. 429-436.

[8] Fordham, P., M. Besson, and P. Gallezot, Selective Catalytic Oxidation of Glyceric

Acid to Tartronic and Hydroxypyruvic Acids. Applied Catalysis A: General, 1995. 133: p.

L179-L184.

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6. Biochemical conversion of glycerol

This chapter studies different possibilities for glycerol bioconversion to added value

products: 1,3-propanediol, ethanol, poly-3-hydroxybutirate, lactic acid, succinic acid,

propionic acid, and rhamnolipids. Also, the influence of the main process variables on the

fermentation behavior (conversion, selectivity, and products distribution) is discussed.

6.1 1,3-propanediol In the early of 90’s a biotechnological route which uses glycerol to produce 1,3-

propanediol by mean a fermentation process was developed [1]. 1,3-propanediol is a

commercially important compound because it can be used as adhesive, antifreeze,

cosmetics moisturizing, stabilizing detergents, and as additive for painting, printing inks,

and high pressure lubricants. Also it can be used as monomer for polyesters synthesis

such as polytrimethylene terephthalate (PTT) and polyethylene terephthalate (PET) which

can improve the chemical and mechanical properties in comparison with other similar

monomers.

Fermentative production of 1,3-propanediol (PD) under anaerobiosis takes place in two

parallel ways. In the first one, a fraction of glycerol is oxidezed by glycerol-dehydrogenase

(Glyc-DH) to dihydroxy-acetone (DHA), and then phosphorrylated by DHA kinase to enter

glycol-lysis. The remaining glycerol is then dehydrated to 3-hydroxypropionaldehyde

(3HPA) by glyceroldehydratase, where reduction con-tinues by

propanedioldehydrogenase (PPD-DH) and by a dependent NAD oxidorreduc-tase to 1,3-

propanediol [2-3]. 1,3-propanediol production can be performed biologically by several

bacterial strains such as Klebsiella pneumoniae, Citrobacter freundii, Enterobacter

agglomerans, Clostridium butyricum, and Clostridium acetobutylicum [4-5]. K.

pneumoniae and C. butyricum are commercially the most promising bacterial strains

because of their high yield, productivity, and resistance to both substrate and product


inhibition. Among these two bacteria, K. pneumoniae DSM-2026 has been presented as

one of the most appropriate bacterial strain for glycerol fermentation to 1,3-propanediol [6]

and it was selected as the main process microorganism in this article. The purpose of this

article is to analyze the glycerol fermentation to 1,3-propanediol by K. pneumoniae in one

and two continuous fermentation stages.

Studies performed under batch and fed-batch cultures have showed low productivities of

1,3-propanediol, about 2-3 g L-1h-1 with a maximum 1,3-propanediol concentration of 50-

60 g L-1. In continuous cultures the productivity can be increased, but the maximum

concentration reached is the half (about 30 g L-1

6.2 Ethanol

) of the obtained under fed-batch or batch

culture conditions. Due to glycerol bioconversion to 1,3-propanediol is a complex

biological mechanism which is subject to inhibitions by substrate and products [7],

process analysis become an important tool to develop efficient configuration process that

allows obtaining the metabolite at high yield, concentration, and productivity. In this sense

Posada et al [8], studied four culture configurations (batch, fed batch, continuous, and two

continuous stages) for 1,3- propanediol production from glycerol, and each configuration

process was optimized.

E. coli has showed the ability for metabolizing glycerol in presence of an external electron

acceptor. Glycerol degradation process begins with the GlpF incorporation in the

cytoplasm. Later phosphorylation process is carried out, which is catalyzed by GlpK

kinasa. This phosphorylated carbohydrate (glycerol 3-phosphate) starts an oxide-

reduction process which is accelerated by different enzymes. The anaerobic process is

catalyzed by the dehydrogenases GlpC, GlpB, and GlpA, while the aerobic process is

catalyzed by GlpD. This dehydrogenation process produces dihydroxyacetone 3-

phosphate and finally glycolysis process takes place to obtain pyruvate (see Figure 6.1.).

Also microorganisms such as Klebsiella, Citrobacter, Enterobacter, Clostridium,

Lactobacillus, Bacillus, Propionibacterium, and Anaerobiospirillum have been reported for

glycerol degrading in fermentative way. Degradation process of these microorganisms is

strongly linked to 1,3-propanediol synthesis with Citrobacter freundii and Klebsiella

pneumonia. However these microorganisms present diverse problems for their industrial

use such as pathogenicity level, requirements of strict anaerobic conditions, and complex

6. Biochemical Conversion 49

cultivation media. In this way, it is necessary to search microorganisms able to metabolize

glycerol without pathogenic effects as occurs with E. coli. Also, E. coli can use glycerol as

carbon source without any external electron receiver. This process is regulated by GldA

dehydrogenase and DHAK dihydroxyacetone kinase for obtaining ethanol, succinate,

acetate, and formate (see Figure 6.2.) [9].

Figure 6.1. Schematic representation of glycerol degradation process on the part of

Escherichia coli, on non fermentative process.

Deletions in E. coli have been carried out to increase formiate and ethanol yields from

glycerol at a concentration of 10 g/L [10]. Thus, from glycerol dehydrogenase (gldA) and

dihydroxyacetone kinase (dhaKLM) over expression a yield of 95% to ethanol from


glycerol was achieved. Also, a genomic analysis was carried out for determining the

genes effect on the change from aerobic to anaerobic conditions in E. coli. A metabolic

characterization to evaluate succinate, acetate, formiate, lactate, and ethanol yields was

carried out [11].

Figure 6.2. Main metabolic pathways for fermentative degradation of glycerol by

Escherichia coli.


A mixture of ethanol and formiate can be produced by glycerol fermentation using

Klebsiella planticola isolated from the rumen [12]. Dharmadi et al [13], reported the

glycerol fermentation by E. coli, the authors have evaluated the pH-dependence and CO2

availability. Ito et al [14], showed that glycerol at 10 g/L was almost completely consumed

within 84 h; the main products were ethanol and succinic acid with molar yields of 86%

and 7%, respectively. According to the authors, E. coli is already a good biocatalyst for

glycerol conversion into ethanol and hydrogen.

E. aerogenes can be used for ethanol production at high yield from biodiesel wastes

containing glycerol. In this way, a synthetic medium containing biodiesel wastes of

glycerol at 80 mM was analyzed and glycerol was consumed in 24 h, producing 0.89 mol

of H2

6.3 poly-3-hydroxybutirate

and 1.0 mol of ethanol per mol of glycerol [14].

Polyhydorxyalcanoates are attractive substitute biopolymers for conventional

petrochemical plastics which have similar physical properties to thermoplastics and

elastomers. PHAs are homo or heteropolyesters synthesized and stored intracellularly by

several Gram Negative bacteria [15]. PHAs can be produced from renewable resources

through a fermentation process under restricted growth conditions for nitrogen,

phosphorus, sulfurs and/or oxygen in the presence of an excess carbon source, and they

can also be completely biodegraded by many microorganisms [16]. PHAs are stored in

form of granules by bacteria and can account for up to 80% of the total bacterial dry

weight [17]. On the other hand, polyhydoroxybutyrates (PHBs) were the first type of PHAs

discovered and the most widely studied. PHB has similar properties to conventional

plastics like polypropylene or polyethylene, and it can be extruded, molded, spun into

fibers, made into films, and used to make heteropolymers with other synthetic polymers

[18-19].

Wild strains such as Cupriavidus necator [20], Methylobacterium rhodesianum or

recombinant microorganism such as E. coli recombinant [21, 22] can produce PHB using

glycerol as a carbon and energy source. Bacteria used for PHAs production can be

divided into two groups based on culture conditions. The first group requires limitation of

an essential nutrient such as N, P Mg, K, O or S, and excess of a carbon source; some


examples are B. megaterium, C. necator, A. eutrophus, P. extorquens, and Ps.

oleovorans. In the second group, nutrient limitation is not required and the polymer can be

accumulated during the growth phase [23]; some examples are E. coli recombinant, Az.

vinelandii recombinant, and A. latus. PHB producer strains which use glycerol as the

carbon source are in the first group of bacteria. Polyhydorxyalcanoates are attractive

substitute biopolymers for conventional petrochemical plastics which have similar physical

properties to thermoplastics and elastomers.

PHAs are stored in the form of granules by bacteria and can account for up to 80% of the

total bacterial dry weight [24]. On the other hand, polyhydorxybutyrates (PHBs) were the

first type of PHAs discovered and the most widely studied. PHB has similar mechanical

properties to conventional plastics like polypropylene or polyethylene, and it can be

extruded, molded, spun into fibers, made into films, and used to make heteropolymers

with other synthetic polymers [25].

The fermentation stage can be performed in different operational modes. Batch PHB

production is normally induced by co-culturing the cells [26] or by limiting them with

nitrogen availability using an excess of carbon source in the stationary phase [27]. To

induce the desired nutrient limitation and to achieve a high cell density, a fed-batch

process is the most commonly used method [28-29]. Thus, cell growth is maintained

without nutrient limitation until a desired concentration is achieved. Then, an essential

nutrient is limited to allow an efficient PHB synthesis. During this nutrient limitation stage

the residual cell concentration (i.e., the difference between cell concentration and polymer

concentration) remains almost constant and cell concentration increases only by

polymeric intracellular accumulation [30]. For bacteria requiring an essential nutrient

limitation a two-stage chemostat should be employed thus resulting in a 1.7 fold higher

productivity compared to the one-stage chemostat [30]. Culture performance is affected

by several variables including temperature, pH, fed carbon-to-nitrogen ratio, concentration

of substrates and trace elements, ionic strength, agitation intensity, and dissolved oxygen.

To substantially enhance the yield and productivity of many bioprocesses, optimization

[31-32] and control [33] of the fermentation conditions has been used.


6.4 D-Lactic acid The exhibited heterofermentative behavior of glycerol metabolism under anaerobic and

microaerobic conditions by wild-type E. coli was recently reported [34-35]. And it was

found that significant amounts of ethanol, acetic acid, succinic acid, and formic acid were

produced, while a negligible amount of D-lactic acid was obtained. Besides the ability of

E. coli to metabolize glycerol under anaerobic and microaerobic conditions, the

corresponding pathways involved in the glycerol utilization were recently elucidated [36]

and under these conditions, ethanol was identified as the primary fermentation product.

Later, based on metabolic engineering strategies, an engineered E. coli for the efficient

conversion of glycerol to D-lactic acid in a minimal medium was reported [37]. Thus, the

homofermentative route to produce D-lactic acid was engineered by overexpressing the

pathways involved in the glycerol conversion to D-lactic acid and blocking the pathways

leading to the synthesis of by-products. In general terms, the enzymes involved in the

pathways for glycerol conversion to glycolytic intermediates (i.e., GlpK-GlpD and GldA-

DHAK) and the enzyme involved in the pathway for D-lactic acid synthesis from pyruvic

acid were overexpressed (i.e., D-lactate dehydrogenase). Meanwhile, the by-products

formation was minimized by inactivation of enzymes such as: pyruvate-formate lyase

(ΔpflB), fumarate reductase (ΔfrdA), phosphate acetyltransferase (Δpta), and

alcohol/acetaldehyde dehydrogenase (ΔadhE). Also, a mutation which blocks the aerobic

D-lactate dehydrogenase (Δdld) was introduced in order to prevent the utilization of D-

lactic acid. The Figure 6.3 shows both the pathways involved in the microaerobic

utilization of glycerol in E. coli and the genetic modifications performed by metabolic

engineering strategies for gene overexpressions or disruptions.

Although lactic acid bacteria have been used for D-lactic acid production from

carbohydrate rich feedstocks, it has also been reported the use of alternative biocatalysts

which are mainly engineered Escherichia coli strains able to produce D- or L-lactic acid

[38-42]. But only a few papers have been published on the use of glycerol as carbon

source for lactic acid production [37, 43]. For instance, Hong et al. [43] compared eight

bacterial strains for lactic acid production from glycerol. Thus, the strain named AC-521

and a member of E. coli, showed the best performance for a fed-batch fermentation

process. On the other hand, Mazumdar et al. [37] engineered several E. coli strains by

overexpressing pathways involved in the conversion of glycerol to lactic acid and blocking


those leading to the synthesis of by-products as it was above described. In all cases they

used a minimal medium supplemented with sodium selenite, Na2HPO4, (NH4)2SO4,

NH4Cl, and 20 (or 40 or 60) g/l of pure (or crude) glycerol.

Figure 6.3. Pathways involved in the microaerobic utilization of glycerol in E. coli and

Genetic modifications supporting the metabolic engineering strategies employed by

Mazumdar et al [37]. Thicker lines (overexpression of gldA-dhaKLM, glpK-glpD, and ldhA)

or cross bars (disruption of pflB, pta, adhE, frdA, and dld). Broken lines illustrate multiple

steps. Relevant reactions are represented by the names of the gene(s) coding for the

enzymes.

6.5 Succinic acid Succinic acid is a C4 dicarboxylic acid produced as both intermediate of the tricarboxylic

acid cycle (TCA) and one of the fermentation products of energy metabolism [44]. This

metabolite can be used for the manufacture of industrially important chemicals including

adipic acid, 1,4-butanediol, tetrahydrofuran, N-methyl pyrrolidinone, 2-pyrrolidinone,

succinate salts and gamma-butyrolactone (see Figure 6.4); and for the synthesis of


biodegradable polymers such as polybutyrate succinate (PBS) and polyamides

(Nylon®x,4) and green solvents [45].

Figure 6.4. Products that can be synthesized from succinic acid.

Succinic acid is currently produced from crude oil by either catalytic hydrogenation of

maleic anhydride to succinic anhydride and subsequent hydration, or direct catalytic

hydrogenation of maleic acid [46]. The commercial price of petrochemically produced

succinic acid is about 5.9–8.8 USD$/kg depending on its purity. Also, for its production

from maleic anhydride, the raw material costs are about 1 USD$/kg of succinic acid [45].

Even though the production of chemicals based on succinic acid accounts to about

16.000 Ton/year [47], the market potential was estimated to be about 270,000 Ton/year if

succinic acid replaced maleic anhydride for all uses [48-49]. Thus, because of these

predictions, the dramatic raising in petroleum price, and the increasing environmental

concerns, the fermentative production of succinic acid from renewable resources has

recently received much attention. In this way several microorganisms including


Actinobacillus succinogenes [50-51], Anaerobiospirillum succiniciproducens [52-53], and

Mannheimia succiniciproducens [54], recombinant Escherichia coli strains [55-56] and

Corynebacterium glutamicum [57-58] have been found to produce succinic. Also, during

fermentative production of succinic acid some by-products such as acetic acid, formic

acid, lactic acid, and ethanol are also obtained. By products formation limits the possibility

of its fermentative production in industrial scale, since the succinic acid yield is reduced

and a more complex and costly downstream process is required [45, 59-60].

The biological production of succinate from glycerol occurs through a redox-balanced

pathway in the presence of excess carbon dioxide. Unlike glycerol, succinate production

from glucose is not redox balanced and can provide a maximum theoretical molar yield of

1.71 (carbon yield of 1.14) without external reducing power. Although ethanol and

succinate are the only two products resulting from redox-balanced pathways of glycerol

fermentation in E. coli [61-62], succinic acid is minor product [61]. In order to improve the

succinate production, Blankschien et al. [63] engineered an E. coli strain by blocking

pathways to competing metabolic products and thus leaving only the succinate pathway

achieving redox balance during glycerol utilization (see Figure 6.5.)

Glycerol dissimilation in E. coli to dihydroxyacetonephosphate (DHAP) can proceed

through two respiratory routes: the aerobic GlpK–GlpD and the anaerobic GlpK–GlpABC,

or through the fermentative route GldA–DhaKLM (see Figure 6.5.). The last one has been

reported to use glycerol efficiently under both anaerobic and microaerobic conditions [61,

64].

Because net ATP is typically not generated by substrate-level phosphorylation when

succinate is produced from glycerol in wild-type E. coli (i.e., through ppc, see below and

Fig. 6.5), use of the fermentative GldA–DhaKLM route is preferred because higher energy

NADH is generated in glycerol dissimilation through GldA as opposed to a reduced flavin

through GlpD or GlpABC [65-66]. However, DhaKLM uses phosphoenolpyruvate (PEP)

as a cofactor, impacting the metabolic nodes available for glycerol fermentation to

succinate (See Figure 6.5).

Production of succinate from glycerol involves fixing CO2 onto a 3-carbon intermediate,

which is stepwise converted to succinate by the reductive branch of the TCA cycle [67]


(See Figure 6.5). E. coli uses PEP carboxylase (ppc) as its main carboxylation enzyme for

succinate generation; however, this is not ideal as PEP levels will be decreased when the

fermentative route of glycerol dissimilation (GldA–DhaKLM) is used (See Figure 6.5). An

analogous argument can be made for the use of the primarily gluconeogenic PEP

carboxykinases (from E. coli or natural succinate producers) [67]. Succinate synthesis

from pyruvate, which is readily available, is limited because E. coli lacks a native

pyruvatecarboxylase (pyc) and the conversion of pyruvate to malate by the gluconeogenic

malic enzymes is not kinetically favored [67]. An effective way to retain the GldA–DhaKLM

route and generate succinate is to introduce a pyruvate carboxylase (pyc) into E. coli,

creating an efficient node for the step wise conversion of pyruvate to succinate.

Figure 6.5. Pathways involved in the micro aerobic utilization of glycerol and the

generation of phosphoenol pyruvate and pyruvate, which can be carboxylated to

intermediates leading to succinate [63].


Use of a heterologous pyruvate carboxylase (pyc) in E. coli to drive succinate production

from glycerol leaves one remaining obstacle, the lack of net ATP production by substrate-

level phosphorylation. Such a complication can be effectively overcome by the use of

microaerobic conditions. ATP will be gained through oxidative phosphorylation resulting

from the reducing equivalents generated during the utilization of glycerol, including those

generated by the incorporation of glycerol into cell mass (i.e. cell mass is less reduced on

average than glycerol) [61] (See Figure 6.5)

6.6 Propionic acid Propionic acid and its calcium, sodium, and potassium salts are widely used as

preservatives in animal feed and human foods, and propionic acid is also an important

chemical intermediate in the synthesis of cellulose fibers, herbicides, perfumes and

pharmaceuticals [68-69]. Currently, almost all propionic acid is produced by chemical

synthesis from petroleum feedstocks. The acid also could be produced by

propionibacteria via the dicarboxylic acid pathway with acetic acid and succinic acid as

byproducts [70-74], but low yield and productivity due to the inhibition of propionic acid on

cell growth and propionic acid synthesis [72, 75] is a problem. To alleviate the inhibition of

propionic acid on microbial growth and propionic acid synthesis, two approaches,

extractive propionic acid fermentation [76-78] and propionic acid production with propionic

acid-tolerant bacteria obtained via adaptive evolution [72, 79-80] have been developed.

Despite such advancements, current microbial propionic acid production cannot

economically compete petrochemical routes. Producing propionic acid from agricultural

and industrial wastes may make microbial propionic acid production economically

competitive. Glycerol is a main by-product of the biodiesel industry [81] and could thus be

a low-cost feedstock to produce propionic acid. While most studies on propionic acid

production by Propionibacterium acidipropionici have focused on glucose and whey

lactose [77, 82-85], some studied have explored glycerol as the carbon source [86-87],

and it was observed that glycerol might be advantageous since less acetic acid was

produced during the consumption of glycerol [70, 86].

Since optimal conditions for the use of glycerol in propionic acid production have not yet

been established, we optimized propionic acid production by propionic acid-tolerant P.

acidipropionici CGMCC 1.2230 with glycerol as the carbon source in batch cultures and


then scaled-up production in a 10 m3

fermentor using the optimized conditions. The

results obtained here may be helpful for industrial production of propionic acid.

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7. Study Cases of Chemical Conversion of Glycerol

Three different technological schemes to transform the glycerol obtained as by-product in

biodiesel industry to added-value products are here designed, simulated, and

economically assessed. Dehydration, steam gasification, and hydrogenolysis were the

analyzed processes where acrolein, hydrogen, and 1,2-propanediol are their respective

products. For dehydration and gasification processes a glycerol conversion of 100% was

reached, and the respective molar yields to acrolein and hydrogen were 85.2 % and 78.2

%. Also, these two processes were heat integrated. 175 and 67 W/(feeding kg) were

recovered for dehydration and gasification respectively. Economic results showed that the

three processes are economically viable, and the highest economical return was obtained

for 1,2-propanediol.

7.1 Generalities Both dehydration and hydrogenolysis reactors were simulated based on a stoichiometric

approach in which acrolein, hydroxyacetone, acetaldehyde, formaldehyde, and water

were considered as the dehydration products, meanwhile 1,2-propanediol,

hydroxyacetone, ethylene glycol, and methane were considered as hydrogenolysis

products. On the other hand, the gasification reactor was simulated as an Rgibbs module

and carbon monoxide, carbon dioxide, hydrogen, methane, and water were considered as

the reaction products.

All processes were integrated in a basic level for a better economic performance. Thus,

the dehydration process was heat integrated, meanwhile the gasification process was

heat and mass integrated and finally the hydrogenolysis process was mass integrated.


In all cases the non-random two liquid (NRTL) thermodynamic model was utilized to

calculate the activity coefficients of the liquid phases and Redlich-Kwong (RK) equation of

state was used to model the vapor phase.

Process engineering looks into the design of high-performance processes that meet

mainly two kinds of criteria: high conversion levels and low production costs. Thus

comparing opportunities to convert the glycerol by-product into added-value products is

the core of this analysis. Also, simulations of the technological schemes were used to

generate their respective mass and energy balance sheets, which are the basic input for

the techno-economic analysis.

7.2 Acrolein Production Acrolein is used as raw material to treat fiber and to produce acrylic acid and medicines,

even more it has been used as a growth control agent of microbes in feed process lines

due to its antimicrobial activity.

Although commercial manufacturing of acrolein has been based on the petrochemical-

propylene oxidation process, this compound can also be produced by homogeneous

catalytic dehydration of glycerol in presence of zinc sulfate. In the last case, for glycerol

dehydration at 360 ºC and 25 MPa, a maximal selectivity of 75 mol % at a conversion

level of 50% was reached [1].On the other hand, both higher selectivity and conversion

were obtained using heteropolyacid catalysts supported on silica with presence of

titanium, aluminium, and zirconium oxides. Thus, the H3PW12O40 catalyst supported on

ZrO2 was able to produce acrolein at a selectivity of 70 % [2]. Also, complete conversion

of glycerol with selectivities ranging from 75 to 86 % was reported at temperatures

between 275 -325 ºC on these heteropolyacid catalysts [3-4]. Even more, silicotungstic

acid [5] and Nb2O5

[6] supported on activated carbon have also been reported to produce

acrolein from glycerol at selectivity levels near to 50%.

In addition to heterogeneous catalysis, acrolein can also be produced by glycerol

conversion on hot-compressed water (HCW) with H2SO4 as catalyst. In this sense, it was

reported that yield to acrolein can be improved by increasing either the operational

7. Study Cases of Chemical Conversion 71

pressure or the concentration of glycerol or H2SO4. Using this conversion way, selectivity

values up to 80% can be obtained [7].

The highly exothermic glycerol dehydration to acrolein is carried out by an acid catalyzed

process as shown in Figure 7.1. An aqueous glycerol stream at 10 wt % is heated in two

stages; in the first one the heat produced during the dehydration reaction is recovered in

the Heat Exchanger I, meanwhile using the Heater I the reaction temperature is reached.

Thus dehydrogenation reaction takes place at 275 ºC and 1 bar.

HE-1 H-1R-1

Con-1

HE-2

DC-1 DC-2

Diluted Glycerol

Reactives

Products Condensate

Vapor

Bottoms 1 Bottoms 2

Distillate 1 Distillate 2

Figure 7.1. Simplified flowsheet for acrolein production by glycerol dehydration.

HE-1: Heat exchanger I; H-1: Heater; R-1: Dehydration reactor; Cond-1: Condenser; HE-

2: Heat exchanger II; DC-1: Distillation column I; DC-2: Distillation column II.

Equations (7.1) to (7.3) describe the reactive system for catalytic glycerol dehydration [4,

8], in which acrolein (C3H4O), hydroxyacetone (C3H6O2), acetaldehyde (C2H4O), and

formaldehyde (CH2

O) are the main reaction products. The normalized yields for each

reaction are 85%, 8% and 7%, respectively.

O2HOHCOHC 243383 +→ (Equation 7.1)

OHOHCOHC 2263383 +→ (Equation 7.2)

OHOCHOHCOHC 2242383 ++→ (Equation 7.3)

After the dehydration reaction, products stream is cooled in the Heat Exchanger I, and

thus this stream is thermally integrated with the fed diluted glycerol stream. Then, a share

of water is condensed and the resulting mixture is cooled to 80 ºC in the Heat Exchanger

II. Thus, the downstream process continues with a distillation column where both

remaining water and hydroxyacetone are retired by the bottoms stream. Finally, to purify

the acrolein stream from 92 to 98.5 wt %, a second distillation column could be used.


Thus, a mixture of acetaldehyde and formaldehyde is obtained by the distillated stream,

but the condenser should be operated with a special coolant liquid since the distillated

stream is obtained to -10 ºC.

During the heating of the fresh feed glycerol stream in the dehydration process, 175

W/(feeding kg) were recovered from the effluent reactor stream by mean of the Heat

Exchanger I. Also, in the dehydration reactor not only glycerol was completely converted

but also a yield to acrolein of 85.25% mol was achieved. Then, following the downstream

process line, acrolein at 92.2 wt % was obtained in the distillated stream from the

Distillation Column I and also 99.4 % of the produced acrolein was recovered.

Additionally, in order to obtain a higher purity of acrolein, a further distillation column was

analyzed. Thus, an acrolein steam at 98.5 wt % was obtained, but to reach the operation

conditions a special coolant is required since the condenser must to operate at -10 ºC. In

this way, when the second distillation column was used, the 98.7 % of the produced

acrolein was recovered, this operational requirements surely increase the production

costs. On the other hand, the most important energy consumptions were obtained the

Heater I and the reboilers of both distillation columns, with net heat duties of: 591.1, 74.4,

and 5.02 W/(feeding kg), respectively. The main simulations results for the dehydration

process are shown in Table 7.1.

Table 7.1. Simulation results for dehydration process from glycerol

Stream

Diluted Glycerol

Reactives

Products

Condensate

Vapor

Distillate 1

Bottoms 1

Distillate 2

Bottoms 2

Temp. (K) 298,1 548,1 548,1 372,6 352,7 309,4 373,1 262,5 325,4 Pressure (atm) 1 1 1 1 1 1 1 1 1 Comp. (wt%) Glycerol 0,1 0,1 0 0 0 0 0 0 0 Water 0,9 0,9 0,936 0,978 0,927 0,006 0,994 0 0,006 Acrolein 0 0 0,052 0,002 0,063 0,922 0 0,087 0,985 Hydroxyacetone 0 0 0,007 0,019 0,004 0 0,004 0 0 Acetaldehyde 0 0 0,003 0 0,004 0,034 0,001 0,371 0,009 Formaldehyde 0 0 0,002 0 0,003 0,038 0 0,542 0 Total Flow Rate (Kg/h) 100 100 100 17,75 82,25 5,59 76,66 0,39 5,20


7.3 Hydrogen production Hydrogen is currently derived from nonrenewable natural gas and petroleum, but it could

be produced from renewable resources such as biomass or its derivates [9]. Many

applications in fields such as electricity generation, fuel cells, and automotive fuels have

been found for hydrogen since it can be used in mobile and stationary applications.

Besides, due to its high energy efficiency, sustainable, and nonpolluting character,

hydrogen is considered as an obvious alternative to hydrocarbon fuels such as gasoline.

Thus, it is expected that hydrogen plays a key role in the world’s energy future by

replacing fossil fuels and storage energy [10].

Non-catalytic processes such as pyrolysis and steam gasification are technologies able to

produce added-value products such as hydrogen and syn gas from glycerol. Pyrolysis is a

thermal cracking process of organic liquids or solids at high temperature performed in

oxygen absence; meanwhile steam gasification is carried out in presence of oxygen and

produces fuel gases with higher hydrogen content than pyrolytic process.

Pyrolysis produces liquid fuels at low temperatures (400 to 600 °C), but when this process

is carried out at high temperatures (> 750 °|C) gaseous products are obtained. Moreover,

gasification is performed in presence of oxygen (i.e., air, pure oxygen, or steam) and a

mixture of carbon monoxide and hydrogen is also produced [11]. In the case of steam

gasification of glycerol at 600 - 700 °C, a yield of 92.3 mol % to syn-gas with a H2

/CO

molar ratio of 2/1 was reported [12]. Meanwhile glycerol pyrolysis over carbonaceous

catalysts at 800 °C produces synthesis gas up to 81 vol % [13].

On the other hand, crude glycerol has been analyzed as raw material to produce

hydrogen. For instance, yields ranging from 77 to 95 wt % were reported for catalytic

steam reforming of crude glycerol on commercial Ni [14]. Besides, higher yields have

been reached by steam gasification from crude glycerol such as 97 % to syn-gas and

65.7 % to H2

[15].

Hydrogen is produced by supercritical water gasification (SCWG), with glycerol as carbon

source. The simplified flowsheet for hydrogen production from glycerol is shown in Figure

7.2. A mixture containing diluted glycerol at 25 wt % is heated in two stages, and an

intermediate compression process is required. Thus, the first heat exchanger produces


overheat vapor and then the reaction pressure is reached by mean of the compressor.

During the compression operation a heat excess is produced, this is used to heat the

fresh diluted glycerol stream in the Heat Exchanger I. Thus, compressed stream and fed

glycerol stream are thermally integrated. Then, by mean of the Hater I the reaction

temperature is achieved, and low heating requirement are needed. The resulting mixture

is fed to the gasification reactor at 600 ºC and 300 bar (i.e. supercritical conditions).

HE-1 H-1R-1

Sep-1

HE-2Diluted Glycerol

Reactives Products 1Vapor 1

Sep-2

Comp-1

Products 2Vapor 2

Condensate 1

Condensate 2

Recycle

Fresh Water

Overheat Vapor

M-1

Figure 7.2. Simplified flowsheet for hydrogen production by gasification.

HE-1: Heat exchanger I; Comp-1: Compressor: H-1: Heater I, R-1: Gasification reactor;

HE-2: Heat exchanger II; Sep-1: High-Pressure (HP) gas-liquid separator; Sep-2: Low-

Pressure (LP) gas-liquid separator; M-1: Mixer.

Molar distribution of reaction products after gasification has been reported as follows [16]:

hydrogen 29.2 %, carbon dioxide 36.1 %, carbon monoxide 0.9 % and methane 33.8 %.

Then, the reaction stoichiometry can be expressed as shows the equation (7.4).

4242383 14CHCOCO15H31OHOH10C +++→+ (Equation 7.4)

Reaction products are cooled in the Heat Exchanger II and a two-phase stream (gas and

liquid phases) is obtained. This mixture is fed to the high-pressure (HP) gas-liquid

separator which operates at temperatures between 25-100 ºC, and at 300 bar. Then the

liquid phase, obtained by bottoms, is further transferred to the low-pressure (LP) gas-

liquid separator which operates at 20 ºC and 1 bar. The HP separator produces a H2-rich

gas stream, while the LP separator produces a CO2-CH4-rich gas. The liquid product from

LP separator, which is mostly water, is mixed with fresh water and fed to the Heat

Exchanger II as cooling fluid, and thus an overheat-vapor stream is obtained. Finally, the

CO2-CH4-rich gas stream could be burned to generate process heat.


For the dehydration process, the feed stream to the gasification process was also heat

integrated. Thus, the heat excess obtained during the isentropic compression operation

(i.e., 67 W/(feeding kg)) was recovered in the Heat Exchanger I. Besides, the gasification

reactor was able to reach a conversion of 100 % at a yield of 78.2 %. In addition to the

heat integration of the fresh feed glycerol stream, the products reaction stream was also

heat integrated and 784 W/(kg of reaction products) were recovered. Thus, an overheat

vapor stream was obtained. In order to purify the hydrogen produced in the gasification

reactor, two gas-liquid separation units were used. The first one is the High-Pressure (HP)

gas-liquid separator in which H2 at 90.9 mol % was obtained, meanwhile in the Low-

Pressure (LP) gas-liquid separator the water was retired and a stream containing a

mixture CO2-CH4 was obtained. Also, this stream containing a mixture CO2-CH4

could be

used to generate process heat. The net heat duties in the gasification process were

represented by the Compressor, Heater I, and Gasification Reactor; where the heat duty

reported by the simulator were 673, 146.3 and -75.13 W/(feeding kg), respectively. The

main simulations results for the gasification process are shown in Table 7.2.

Table 7.2. Simulation results for gasification process from glycerol

Stream

Diluted

Glycerol

Reactives

Product

s

1

Products

2

Vapor

1

Condensate

1

Vapor

2

Condensat

e 2

Recycle

Temp. (K) 298,1 873,1 873,1 308,9 308,9 308,9 293,1 293,1 293,1

Pressure (atm) 1 300 300 300 300 300 1 1 1

Comp. (wt%)

Glycerol 0,061 0,061 0 0 0 0 0 0 0

Water 0,939 0,939 0,804 0,804 0 0,841 0,023 0,978 0,984

CO 0 2 0 0,074 0,074 0,032 0,076 0,451 0,013 0,009

CO 0 0 0,002 0,002 0,007 0,002 0,011 0 0

CH 0 4 0 0,080 0,080 0,052 0,082 0,515 0,009 0,006

H 0 2 0 0,041 0,041 0,909 0 0 0 0

Total Flow Rate (Kg/h) 100 100 100 100 0,994 99,01 20,69 78,32 107,99


7.4 1,2-propanediol production 1,2-propanediol is produced commercially by the the hydration of propylene oxide derived

from propylene by either the chlorohydrin process or the hydroperoxide process. But in

the presence of metallic catalysts and hydrogen, glycerol can be hydrogenated to

propylene glycol, 1,3 propanediol, or ethylene glycol.

1,2-propanediol is used in unsaturated polyester resins, functional fluids (antifreeze, de-

icing, and heat transfer), pharmaceuticals, foods, cosmetics, liquid detergents, tobacco

humectants, flavors and fragrances, personal care, paints and animal feed [17]. The

antifreeze and deicing market is growing because of concern over the toxicity of ethylene

glycol-based products.

1,2-propanediol is obtained from glycerol by the sequential processes of

dehydrogenation-hydrogenation via hydroxiacetone, also glycerol hydrogenolysis can be

carried out in both liquid and vapor phases. Liquid process takes place at low pressure

producing mainly 1,2-propanediol and 1,3-propanediol in presence of supported catalysts

such as Rh [18], Ru [19], or Pt [20]. On the other hand, glycerol hydrogenolysis in vapor

phase is catalyzed by Cu at high hydrogen pressure [21], but in this process lateral

reactions occur and different reaction by-products are obtained. In order to overcome this

problem a two-step process has been proposed, thus dehydrogenation is first performed

under vacuum conditions and then hydrogenation is carried out at high hydrogen pressure

[22-23].

An efficient two steps process for selective production of 1,2-propanediol from glycerol

was developed [24]. The reaction is carried out in vapor phase on a copper metallic

catalyst at ambient pressure of hydrogen. In the first step hydroxyacetone is produced by

glycerol dehydrogenation and then 1,2-propanediol is produced by hydrogenation of

hydroxyacetone. Also, due to both the high temperature required for the dehydrogenation

reaction and the improved selectivity at low temperatures for the hydrogenation reaction,

the reactor configuration has a gradient temperature ranging from 200 °C (on the reactor

top) to 120 °C (on the reactor bottom). Thus, a molar selectivity of 96% to 1,2-propanediol

with total glycerol conversion is achieved.


The simplified flowsheet for 1,2-propanediol production from glycerol is shown in Figure

7.3. Glycerol is diluted at 30 g/L and then this stream is heated up to the reaction

temperature. This stream is fed to the dehydrogenation-hydrogenation reactor, besides a

hydrogen stream at 200 °C is fed in a volumetric ratio of 1/141 (glycerol/H2

) to the reactor.

The gradient reactor is available to convert glycerol completely to 1,2-propanediol,

hydroxyacetone, ethylene glycol, and methane, as is shown by the equation system (7.5)

to (7.7).

DC-1

Glycerol

Bottoms 1

Distillate 1

DC-1

Bottoms 2

Distillate 2

Water

M-1 H-1R-1

D-1

M-2 H-2

HE-1Sep-1

Reactives Products

Condensate

Purge

Recycled H2

Reactive H2 Fresh

H2

Figure 7.3. Simplified flowsheet for 1,2-propanediol production by hydrogenolysis.

M-1: Mixer I; H-1: Heater I; R-1: Hydrogenolysis reactor; HE-1: Heat exchanger; Sep-1:

Gas-líquid separator; D-1: Divisor; DC-1: Distillation column I; DC-2: Distillation column II;

M-2: Mixer II; H-2: Heater III.

OHOHCOHC 2263383 +→ (Equation 7.5)

2832263 HOHC OHC→+ (Equation 7.6)

42622263 H2OHC CHOHC +→+ (Equation 7.7)

1,2-propanodiol purification process starts with a flash operation at 30 °C and 50 bar.

Then, the gas stream containing mainly H2 is recycled and mixed with fresh H2. The

resultant stream of H2 must be heated up to reaction temperature and then fed to the

reactor. The liquid stream obtained from the flash operation is purified using two

distillation columns. In the first one, most of the water quantity is retired meanwhile in the

second distillation column the remaining water and the non-converted hydroxyacetone are

retired.


During the simulation of the hydrogenolysis process, the fresh feed glycerol stream was

heated up to 200 °C. Then, a hydrogen stream at the reaction temperature was fed to the

reactor at a volumetric ratio of 1/141 (glycerol/H2) and glycerol is completely converted in

the hydrogenolysis reactor. The purification process for 1,2-propanediol requires an

evaporation process at 30 °C and 50 bar, in which a gaseous stream containing mainly

H2 (99.9 mol %) was recycled and mixed with fresh H2

. The liquid stream obtained by the

bottom stream from the evaporation process was purified by mean of two distillation

columns. Thus, 99.96 % of water was discarded using the Distillation Column I, and both

the remaining water and no-converted hydroxyacetone were obtained by the distillated

stream in the Distillation Column II. In this way, 1,2-propanediol at 99 wt % was achieved.

The main simulation results for the hydrogenolysis process are shown in Table 7.3.

Table 7.3. Simulation results for hydrogenolysis process from glycerol

Stream

Reactives Products Condensate Bottoms 1 Bottoms 2 Recycled H2 Reactive H

Temp. (K) 2

200 200 30 183.6 186.7 30 200

Pressure (bar) 1 1 50 1 1 50 1

Comp. (wt%)

Glycerol 0,3 0 0 0 0 0 0

1,2-PD 0 0,020 0,240 0,969 0,99 0 0

Hydroxyacetone 0 0,001 0,007 0,029 0,01 0 0

Ethylene glycol 0 0 0 0,001 0,001 0 0

Water 0,7 0,069 0,753 0,001 0 0,008 0,007

H2 0 0,911 0 0 0 0,992 0,992

Total Flow Rate

(Kg/h) 33333 408099 33398 8263 7997 355966 374765

7.5 Economic assessment Operational and capital costs were disaggregated by raw material, utilities, operation

labor, maintenance, operating charges, plant overhead, general and administration costs,

and depreciation cost, as shows the Table 7.4 for each glycerol conversion process [25-

32]. Besides, all costs were normalized and their shares are shown in Table 7.5.


Table 7.4. Production costs for glycerol conversion to added-value

Production costs (US$/L of product)

Acrolein at 92 wt %

Acrolein at 98.5 wt %

Hydrogen

1,2-propanediol

Raw materials 0,2927 0,2920 4,777E-05 0,1203

Utilities 0,3067 1,2006 3,731E-05 0,0813

Operating labor 0,0033 0,0042 1,129E-05 0,0073

Maintenance 0,0116 0,0123 1,031E-05 0,0066

Operating charges 0,0008 0,0011 2,823E-06 0,0081

Plant Overhead 0,0069 0,0082 1,080E-05 0,0063

General and Administrative

0,1499

0,2105

9,625E-06

0,0191

Depreciation of capital 0,0555 0,0646 1,069E-4 0,0192

Total costs 0,8274 1,7935 2,368E-4 0,2682

Sale Price 1,110 1,779 2,498E-4 0,4200

Table 7.5. Percentage of Production costs for glycerol conversion to added-value

Production costs (US$/L of product)

Acrolein at 92%

Acrolein at 98.5%

Hydrogen

1,2-propanediol

Raw materials 35.38 16.28 20.17 44.85

Utilities 37.07 66.94 15.75 30.31

Operating labor 0.40 0.23 4.77 2.72

Maintenance 1.40 0.69 4.35 2.46

Operating charges 0.10 0.06 1.19 3.02

Plant Overhead 0.83 0.46 4.56 2.35

General and Administration 18.12 11.74 4.06 7.12

Depreciation of capital 6.71 3.60 45.14 7.16

Although glycerol at technical grade was considered as raw material for the three

technological schemes, some differences in the raw material costs can be observed due

to the differences among yield, selectivity, and products density (see Table 7.4). For


instance, because of the dehydration process had the lowest selectivity; the highest raw

material cost was obtained for acrolein. Also, due to hydrogen is only product here

obtained in gas phase not only the lowest raw material cost but also the lowest production

cost in US$/L were obtained for the gasification process. Then, in order to compare the

three technological schemes, it was necessary to include the commercial sale price for

each product as shown Table 4. Even more, the shares of each item allow identifying the

main economical resource consumers, as shown Table 7.5.

Two qualities of acrolein are observed in Table 7.4, there are 92 and 98.5 wt %. These

assessment were performed because of the acrolein production process at 98.5 wt %

requires a powerful coolant system which implies high operational costs; and thus its total

production cost is higher than the commercial sale price. On the other hand, since the

service cost to produce acrolein at 92 wt % is only the 25 % required to obtain acrolein of

high purity, the total production cost for this process is lower than its commercial sale

price. During the acrolein production at 92 wt %, most of the production costs are

represented by raw materials and services which totaling 72 % of the total production

cost. Meanwhile for acrolein production at 98.5 wt %, only the services contributing the 67

% of total production cost. Thus, the production process of acrolein at high purity is not

economically viable.

Although in most of the chemical process the raw material cost represents near to 50 % of

the total production cost, in the case of hydrogen production from glycerol the sum of both

raw material and services costs were almost 36 % of the total production cost. But, for this

process the main investment is represented by the process units since the equipment

depreciation is 45.14% of the total production cost. The high depreciation cost occurs

because extreme operational conditions (i.e., high temperatures and pressures) are

required during this process.

Finally for 1,2-propanediol production the raw material and services costs represent the

most share of the total production cost, which add 75%. And then, the equipment

depreciation cost is 16.7 % of the total production cost. These values are typical for a

chemical process.


By comparing the ratio of the commercial sale price respect to the obtained total

production cost (i.e., sale/production costs), a ratio of 1.055 was found for hydrogen

production, followed by the acrolein process with a ratio of 1.34, and the highest value

obtained was for 1,2-propanediol production at a ratio of 1.57. Thus, the production of 1,2-

propanediol could generate the highest economical return for glycerol conversion into

added-value products among the analyzed processes.

7.6 Conclusions Acrolein, hydrogen, and 1,2-propanediol, are three of the most commercially important

products obtained from glycerol, due to their applications, established market, and sale

prices. Here the technological schemes to produce these compounds were designed,

simulated, and economically assessed. Thus, simulation results showed that all the

processes are technologically feasible reaching high purity of product. Also, acrolein

production was found to be viable at a purity of 92 wt %, but do not at a purity of 98.5 wt

%. Finally, both hydrogen and 1,2-propanediol production processes are also

economically viable, where the last one generates the highest profit margin.

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8. Study Cases of Biochemical Conversion of Glycerol

This chapter presents the process design, simulation, and economical assessment of six

different possibilities for glycerol transformation by fermentation. For the production of 1,3-

propanediol, a kinetic model was used allowing optimizing the fermentation stage by three

different approaches. For ethanol, poly-3-hydroxybutirate, lactic acid, succinic acid, and

propionic acid production, a yield approach was used in all cases. Thus, several

scenarios were analyzed in each case depending on the glycerol fermentation stage or on

the downstream process.

8.1 1,3-Propanediol production Although 1,3-propanediol could be biologically produced from glycerol by several bacterial

strains such as: Klebsiella pneumoniae, Citrobacter freundii, Enterobacter agglomerans,

Clostridium butyricum, and Clostridium acetobutylicum [1-2]; the K. pneumoniae and C.

butyricum strains are the most promising bacterial because of their high yield,

productivity, and resistance to both substrate and product inhibition. Among these two

bacteria, K. pneumoniae DSM-2026 has been presented as one of the most appropriate

bacterial strain for glycerol fermentation to 1,3-propanediol [3]. First of all, here the

fermentation process of glycerol to 1,3-propanediol by K. pneumoniae is analyzed in one

and two continuous stages.

The material balances for continuous glycerol fermentation in one single stage are solved

using two independent variables namely the glycerol concentration in the feed stream and

the dilution rate. Since, acetic acid and ethanol are also produced during glycerol

fermentation to 1,3-propanediol by K. pneumoniae, the material balances in a dynamic

state needed to be solved for biomass, substrate, and all the obtained products as shown

in equations (8.1) to (8.5).


Outii

Outi

Inii

Outi XXXD

dtdX

µ ) ( +−= (8.1)

OutiiG

OutiG

IniGi

OutiG XqCCD

dtdC

,,,, ) ( −−=

(8.2)

OutiiPD

OutiPD

IniPDi

OutiPD XqCCD

dtdC

,,,, ) ( +−=

(8.3)

OutiiHAc

OutiHAc

IniHAci

OutiHAc XqCCD

dtdC

,,,, ) ( +−=

(8.4)

OutiiEtOH

OutiEtOH

IniEtOHi

OutiEtOH XqCCD

dtdC

,,,, ) ( +−=

(8.5)

Where: X, CG, CPD, CHAc, and CEtOH, are the concentrations for biomass (g/L), glycerol

(mol/L), 1,3-propanediol (mol/L), acetic acid (mol/L), and ethanol (mol/L), respectively. D

is the dilution rate (i.e., ratio between volumetric flow and reactor volume) (h-1), μ is the

specific rate of cellular growth (h-1), qG is the specific rate of glycerol consumption, and

qPD, qHAc, and qEtOH are the generation rates of each product (h-1

). Subscript i, indicates

the fermentation stage for a multistage system. In the i fermentation stage, In and Out

superscripts indicate the in and the out conditions respectively.

The kinetic model of glycerol fermentation by K. pneumoniae has been previously

explained [4-6]. Specific rates of cell growth, substrate consumption, and products

formation are given in the equations (8.6) to (8.11).

−

−

−

−

+=

∗∗∗∗EtOH

iEtOH

HAc

iHAc

PD

iPD

G

iG

SiG

iGi C

CCC

CC

CC

KCC ,,,,

,

,max 1111µµ

(8.6)

∗+∆++=

SiG

iGmGm

G

iGiG KC

Cq

Ymq

,

,,

µ

(8.7)

∗+∆++=

PDiG

iGmPD

mPDiPDiPD KC

CqYmq

,

,, *µ

(8.8)

∗+∆++=

HAciG

iGmHAc

mHAciHAciHAc KC

CqYmq

,

,, *µ

(8.9) m

iGEtOHiGiEtOH Yqq ),/(,, = (8.10)

iGiiGi

miGEtOH CDc

bCDc

bY,2

2

,1

1),/( +

++

= (8.11)

8. Study Cases of Biochemical Conversion 87

Equation (8.6) describes the specific rate of cell growth which represents a kinetic model

with inhibition by both substrate and products. The kinetic parameters , , ***HAcPDG CCC and

*EtOHC are the critical concentrations (i.e., the concentration where biological activity is

stopped). The required parameters to solve the kinetic model (valid at 37 ºC and at

neutral pH [4, 6]) are: Maximum Specific Growth Rate, μmax=0.67 h-1

and the Monod

Saturation Constant, Ks=2.8e-4 gmol/L.

Constants for specific rate of substrate consumption and product formation are: mG = 2.20

e-3, mPD=-2.69e-3, mHAc=-9.7e-4, YmGly= 8.2, Ym

PD=6.769e-2, YmHAc=3.307e-2, Δqm

Gly

=2.858e-2, ΔqmPD=2.659e-2, Δqm

HAc=5.74e-3, K*Gly=1.143e-2, K*PD=1.55e-2, K*HAc=

8.571e-2. Also, b1, b2, c1, and c2

constants in equation (8.11) are: 2.5e-5, 5.18e-3, 6e-5,

and 5.045e-2 mol/(L*h), respectively.

The critical concentrations required in equation (8.6) were taken from [2]. These

concentrations have an average global deviation of 8.6% for 29 stable states of glycerol

fermentation by K. pneumoniae and C. butyricum [5-6]. The critical concentrations for

glycerol, 1,3-propanediol, acetic acid, and ethanol are: 2.012, 0.8975, 0.7798, and 0.3975

mol/L, respectively.

On the other hand, glycerol fermentation to 1,3-propanediol by K. pneumoniae presents a

metabolic overflow of products causing dynamic phenomena of non-lineal behavior such

as multiplicity of steady states, hysteresis, and oscillations [5]. Thus, the operational

conditions that take to multiplicity of steady states were determined.

In order to obtain the best performance in the first fermentation stage, the volumetric

productivity was optimized using the Levenverg-Marquardt method [7]. Volumetric

productivity is one of the most important functions to be optimized from an operative point

of view, since it implies a high production in a small reactive volume on a short period of

time. The volumetric productivity is shown in equation (8.12), where Pr1 is the volumetric

productivity in (mol/(L*h)), CPD1 is the outlet 1,3-propanediol concentration in (mol/L),and

D1 is the dilution rate in (h-1

). The first fermentation stage in this equation is indicated by

the subscript 1.


1PD11 DCPr = (8.12)

In this case, when multiplicity occurs, the steady states with the highest concentration of

1,3-propanediol were considered. This is to ensure that the steady states with the higher

volumetric productivities are evaluated. Regardless of the inhibition phenomenon caused

by substrate or products, fermentation systems could be successfully performed in two

continuous stages reaching simultaneously a high productivity and a high product

concentration in the first and second fermentation stages respectively [4, 8].

To assess the glycerol fermentation process in two continuous stages, three optimization

models were analyzed. The first optimization model is a sequential procedure in which the

outlet stream from the first fermentation stage (operated at optimal conditions) is directly

fed on the second one. The volumetric productivity on the second fermentation stage was

calculated as a function of both the dilution rate and the achieved change on the 1,3-

propanediol concentration as shown in equation (8.13). Also, the maximum outlet

concentration of 1,3-propanediol on the second fermentation stage was determined by the

Levenverg-Marquardt method.

) ( Pr 1222 PDPD CCD −= (8.13)

Although in the second optimization model the objective function is the volumetric

productivity in the second fermentation stage (see equation 8.13), only the optimal dilution

rate in the first fermentation stage was kept unchanged. Thus, the volumetric productivity

was calculated as a function of both the dilution rate on the second stage and the feed

glycerol concentration on the first stage.

Finally, in the third optimization model the productivity of both fermentation stages were

simultaneously considered as a function of the dilution rate in both stages and the feed

concentration of glycerol in the first fermentation stage. The objective function in this

model is the product of productivities between both fermentation stages (Pr3

), as shown in

equation (8.14).

213 PrPrPr = (8.14)


Figure 8.1 shows the multiple steady states, hysteresis loops, and the wash out line for

the continuous glycerol fermentation. Multiplicity of steady states and hysteresis loops

were studied considering the dilution rate as a parameter from a low glycerol

concentration up to the wash out conditions. In the hysteresis loops when the feed

glycerol concentration increases, low 1,3-propanediol yields are obtained. Moreover,

when the feed glycerol concentration decreases, high 1,3-propanediol yields are acquired.

The latter condition corresponds to the upper curves in the hysteresis loops. The "wash

out" line indicates the extreme operational conditions where the dilution rate equals to the

cellular growth rate ( ii D=µ ).

Thus, volumetric productivity was calculated as a function of both the feed glycerol

concentration and the dilution rate, using a polygon mesh (partition) from low feed

glycerol concentration up to the wash out line (see Figure 8.2.a.). The conditions that

generate the higher 1,3-propanediol concentration were selected at the multiplicity of

steady states region. As a consequence, a discontinuity between A and B was obtained

and the higher volumetric productivity values were located at the right side of these

points.

Figure 8.1. Hysteresis loops and multiple steady states. Concentrations of: a) Biomass

(g/L). b) Residual Glycerol (mol/L). c) 1,3-Propanediol (mol/L). d) Acetic Acid (mol/L) and

e) Ethanol (mol/L). Vertical lines indicate the limits of the multiple steady states region.

Dotted lines show the "wash out" conditions for each dilution rate.


Figure 8.2.a) 1,3-propanediol volumetric productivity, (the column in the right side gives

the scale). b) Region of multiplicity of steady states, optimal productivity for each dilution

rate, global optimal productivity, and wash-out line.

For each dilution rate exists a feed glycerol concentration that generates a maximum

volumetric productivity as obtained by the polygon mesh distribution shown in Figure 2.a.

In order to obtain the global optimum for volumetric productivity both independent

variables (i.e., feed glycerol concentrations and dilution rate) must be simultaneously

considered. Since the highest volumetric productivity is close the steady states region,

this area must be considered when selecting the initial estimated to apply in the

optimization method.

In order to find the conditions for feed glycerol concentration and the dilution rate that

generates the highest volumetric productivity, the Levenverg-Marquardt optimization

method was employed [7]. Since volumetric productivity is a non-continuous function, the

initial estimated for both the feed glycerol concentration and the dilution rate must be

higher than the conditions obtained in the A point.

The multiple steady states region for glycerol fermentation was reported by Xiu et al. [26],

but the used critical concentration parameters have a smaller fitting than the ones used in

this work. The optimal conditions for glycerol fermentation in one continuous stage are as

follows: 0.2821 h-1

for the dilution rate and 0.6882 mol/L for the feed glycerol

concentration, with a volumetric productivity of 0.1076 mol/(L*h).


Additionally, the obtained outlet concentration of 1,3-propanediol is 0.3811 mol/L. This

optimal volumetric productivity is outside the multiple steady states region as shown in

Figure 8.2.b. Since this volumetric productivity is very close to the multiple steady states

region, minimum requirements in the automatic control of the equipments are

recommended. Additionally, Figure 8.2.b. shows the wash out conditions and the optimal

volumetric productivity for each dilution rate.

The kinetic model given for the fermentation system by equations (8.1) to (8.12) is equally

applicable for simulation of a fermentation process with two continuous stages. Thus,

three models to optimize the second fermentation stage were used.

In the first model, optimal conditions for the first fermentation stage were used to calculate

the volumetric productivity on the second fermentation stage, but this function increases

proportionally to the dilution rate, contrary to the behavior shown by the concentration of

1,3-propanediol (see Figure 8.3). The optimal concentration of 1,3-propanediol was

0.4126 mol/L at a dilution rate of 1.9850 h-1

, with a productivity of 0.0625 mol/(L*h).

Figure 8.3. 1,3-Propanediol productivity and concentration in the second fermentation

stage.


In the second optimization model the dilution rate for the first stage was kept from the first

model and the productivity in the second fermentation stage was optimized as a function

of the feed glycerol concentration. The reached productivity is 0.1128 mol/(L*h), at a

dilution rate on the second stage of 0.79 h-1

, with a feed glycerol concentration on the first

stage of 0.8817 mol/L. The outlet concentration of 1,3-propanediol from the first and

second stages are 0.3405, and 0.4833 mol/L, respectively. Since each dilution rate has its

own optimal productivity, it was necessary to calculate the global optimal productivity

using the Levenverg-Marquardt method (see Figure 8.4). Optimal productivities for each

dilution rate in the second fermentation stage are represented by the discontinuous curve.

Also, the P point indicates the global optimal productivity in the second stage.

Figure 8.4. Volumetric productivity in the second fermentation stage using the optimal

dilution rate obtained by the model 1 for the first fermentation stage, (the column in the

right side gives the scale).

Finally, in the third optimization model the productivities of both fermentation stages were

simultaneously optimized considering the two dilution rates and the feed glycerol

concentration as independent variables. The obtained results using the optimal dilution

rate in the first fermentation stage are shown in Figure 8.5. The optimum product of

productivities for each dilution rate in the second fermentation stage is represented by the

discontinuous curve.


Figure 8.5. Product of productivities of both fermentation stages using the optimal dilution

rate obtained by the model 1 for the first fermentation stage, (the column in the right side

gives the scale).

The Q point indicates the global optimum for the product of productivities. This point

corresponds to a dilution rate of 0.2821 h-1 and 3.08 h-1 in the first and second stages

respectively, and a feed glycerol concentration of 0.7362 mol/L in the first stage. The

reached product of productivities was 0.0116 (mol/(L*h))2

where the outlet concentration

of 1,3-propanediol was 0.4124 mol/L and the global molar yield was 0.5602 1,3-PD/ Gly.

Table 8.1 shows the results of the three different models used to optimize the

fermentation of glycerol in two stages. The highest global yield and volumetric productivity

in the first fermentation stage were generated using the sequential optimization model. On

the other hand, the volumetric productivity in the second fermentation stage was

optimized using the combined optimization model under the optimal dilution rate in the

first fermentation stage. But, when the dilution rate in the first fermentation stage was

decreased at 0.25 h-1, the best final concentration of 1,3-propanediol was obtained as

shown in Table 8.1.


Table 8.1. Results summary for each optimization model.

Optimization Model

CS0

a D1

b D2

c CPD1

d CPD2

e YPD/S

f Pr1

g Pr2

h Pr3

i

Sequential. 0,688 0,282 1,985 0,381 0,413 0,599 0,107 0,063 6,72e-3

Combined 0,882 0,282 0,790 0,341 0,483 0,548 0,096 0,113 1,09e-2

Combined 0,932 0,250 0,940 0,396 0,512 0,549 0,099 0,109 1,08e-2

Combined 0,852 0,300 0,720 0,311 0,469 0,551 0,093 0,114 1,06e-2

Simultaneous 0,736 0,282 3,080 0,377 0,412 0,560 0,106 0,109 1,15e-2 a Feed glycerol concentration in the first stage in (mol/L), b Dilution rate in the first

fermentation stage in (h-1), c Dilution rate in the second fermentation stage in (h-1), d 1,3-

propanediol concentration in the first fermentation stage in (mol/L), e 1,3-propanediol

concentration in the second fermentation stage in (mol/L), f Global fermentation yield, g

Volumetric productivity in the first fermentation stage in (mol/(L*h)), h Volumetric

productivity in the second fermentation stage in (mol/(L*h)), i Product of productivities in

(mol/(L*h))2

.

Also, when the dilution rate in the first fermentation stage was increased to 0.30 h-1

, the

highest volumetric productivity in the second stage was obtained as shown in Table 8.1.

Then, the volumetric productivity in the second fermentation stage can be optimized only

at a specific dilution rate in the first fermentation stage.

Using the simultaneous optimization model a high volumetric productivity in both

fermentation stages and the highest product of productivities were obtained. Also, the

obtained value for the optimal dilution rate in the first stage was the same using the

sequential optimization model. Thus, the use of a sequential optimization model allowed

obtaining the highest global yield for 1,3-propanediol (0.599) and the maximum volumetric

productivity in the first fermentation stage (0.1075 mol/(L*h)), whereas the highest 1,3-

propanediol outlet concentration (0.512 mol/L) was observed when the combined

optimization model was employed. Meanwhile, using the simultaneous optimization model

showed both: high volumetric productivities in the two fermentation stages and the highest

product of productivities (0.01157 (mol/(L*h))2). In this way, for the fermentation of

glycerol in two continuous stages, three different operational configurations are available

depending on the desired process objective namely global yield, 1,3-propanediol outlet

concentration, or high simultaneous productivity.


On the other hand, the main problem for designing the downstream process for 1,3-

propanediol recovery and purification from the fermentation broth is the high hydrophilicity

and high boiling point of 1,3-propanediol. The purification scheme proposed here is based

on integrated reaction-separation units to carry out the recovery of 1,3-propanediol. The

first integrated stage is a reactor-extraction process where the hydrophilic nature of 1,3-

propanediol is changed by the acetylation reaction with iso-butyl aldehyde, which

produces 2-iso-propyl-1,3-dioxane as shown in Figure 8.6.

OHOH +CH3

CH3 O CH2

CH2

CH2

O

CH

OCH3

CH3

OH2+H+

Figure 8.6. Acetylation reaction of 1,3-propanediol with iso-butyl aldehyde to 2-iso-propyl-

1,3-dioxane

The 2-iso-propyl-1,3-dioxane has a hydrophobic character which is dragged into the

organic phase containing mainly iso-butyl aldehyde. This aldehyde acts as both reagent

and solvent for the reactive-extraction process [9-11]. Subsequently, the 1,3-propanediol

is recovered by reactive distillation of 2-iso-propyl-1,3-dioxane and water though out the

reverse reaction of cyclical acetylation. Thus, 1,3-propanediol at high purity is obtained by

bottom while the iso-butyl aldehyde is recovered by distillated and it is able to be reused

in the downstream process. Then, the aldehyde is not consumed in the purification

process.

Based on the fermentation results obtained previously for the fermentation stage, three

scenarios are selected in order to perform the process design and analysis for glycerol

conversion to 1,3-propanodiol.

The first scenario considers conditions of optimal volumetric productivity in the first

fermentation stage and optimal final concentration of 1,3-propanediol according to the

sequential model presents in Table 8.1. The second scenario considers conditions of the

highest final concentration of 1,3-propanediol and the highest productivity in the second

fermentation stage according to the second combined model presented in Table 8.1. And

finally, the third scenario considers the optimal global productivity having into account

both fermentation stages according to simultaneous model presented in Table 8.1.


Calculated results of the two-stage fermentation process for each scenario are shown in

Table 8.2, where the column named Feed indicates the fed stream to the first

fermentation stage, Products 1 corresponds to the fermentation products stream from the

first fermentation stage, and Products 2 is the fermentation products stream from the

second fermentation stage.

The normalized stoichiometry for the fermentative reactions are shown in Table 8.3

according to fermentation results obtained for each scenario (Scen. in Table 8.3) and

each fermentation tank (Ferm. in Table 8.3). Also, the molecular formula used for K.

pneumoniae is CH1.75O0.46N0.23

.

Table 8.2. Fermentation results for the three considered scenarios

Concentration mmol/L Feed Products 1 Products 2

Scenario 1

Glycerol 688.2 77.6 8.9

13PD 0.0 381.1 419.7

AcAc 0.0 109.6 122.5

EtOH 0.0 43.4 51.3

Biomass (g/L) 0.0 2.7741 3.1788

Scenario 2

Glycerol 932.0 330.1 153.6

13PD 0.0 395.1 511.4

AcAc 0.0 118.2 149

EtOH 0.0 23.6 35.1

Biomass (g/L) 0.0 2.4924 3.1489

Scenario 3

Glycerol 736.2 147.8 92.1

13PD 0.0 376.8 412.2

AcAc 0.0 111.7 121.7

EtOH 0.0 33.4 37.3

Biomass (g/L) 0.0 2.6267 2.8725


Table 8.3. stoichiometric reactions for each scenario and each fermentation stage

Reactions Glycerol Residual Gly 13PD AcAc EtOH Biomass Molecular Weight 92.09 92.09 76.09 60.05 46.07 23.94

Scen 1, Ferm 1 1.0 0.1067 0.5238 0.1506 0.0597 0.1593

Scen 1, Ferm 2 1.0 0.1045 0.4530 0.1514 0.0927 0.1984

Scen 2, Ferm 1 1.0 0.3399 0.4069 0.1217 0.0243 0.1072

Scen 2, Ferm 2 1.0 0.4523 0.3424 0.0907 0.0339 0.0807

Scen 3, Ferm 1 1.0 0.1896 0.4834 0.1433 0.0429 0.1408

Scen 3, Ferm 2 1.0 0.4832 0.3071 0.0868 0.0338 0.0891

The simplified flowsheet for 1,3-propanediol production from raw glycerol is shown in

Figure 8.7. Cell mass contained in the fermentation is withdrawn throughout a

centrifugation process. Then, the clarified fermentation broth is mixed with iso-butyl

aldehyde in a weight ratio of 5/1 (iso-butyl aldehyde /1,3-propanediol). The reactive-

extraction process takes place at 10 °C since better distribution coefficients are obtained

at lower temperatures. From the reaction extraction unit two product streams are

obtained.

Aqueous stream should be purified in order to recover significant amounts of both iso-

butyl aldehyde and 2-iso-propyl-1,3-dioxane. This stream is subjected to two distillation

stages, in the first one the heterogeneous azeotrope iso-butyl aldehyde-water is obtained

by distillated, and then iso-butyl aldehyde is obtained at 97.3 wt % by decantation. In the

second distillation column the homogeneous azeotrope composed by 2-iso-propyl-1,3-

dioxane and water is obtained at the top of the distillation column. Then, the obtained

azeotrope is mixed with the organic stream obtained during the reactive extraction

process which contains mainly iso-butyl aldehyde and 2-iso-propyl-1,3-dioxane. The

resulting mixture is distilled and the heterogeneous azeotrope iso-butyl aldehyde-water is

once again obtained as the top product, and thus a mixture containing 2-iso-propyl-1,3-

dioxane and water is obtained. This stream is directly fed to the reactive distillation

column as follows.


Raw Glycerol

Metanol

Solids

Water

OrganicPhase

Aqueous glycerol

Waste Water 1 Waste

Water 2

Glycerol(85wt%)

E-1R-1 C-1 Dec-1

E-2

DC-1RII-1

Glycerol(98wt%)

Adsorbate

Solids

Water

Bottoms-2

Diluted Glycerol Fermentation

Broth

Distillated-2

M-1

F-1

Cen-1DC-2

Fresh IBuAld

RE-1

Aqueous Phase

M-2

Bottoms-3

Distillated-3 DC-3

Waste water

Bottoms-4

Distillated-4

DC-4

13PD

Distillated-6

DC-5

Distillated-5

RDC

Water

OrganicPhase

M-2

Dec-2

Bottoms-5

Waste water

IBuAld

Dec-3

F-2

P-3

Figure 8.7. Simplified flowsheet for 1,3-propanediol production from raw glycerol. E:

evaporator, R: reactor, C: centrifuge, Dec: Decanter, DC: distillation column, M: Mixer, F:

fermentator, RE: Reactor extractor, RDC: Reactive distillation column.

In order to determine both the operational viability and the best configuration in the

reactive-distillation tower (localization of the reaction zone), the Static Analysis is applied

[12-16]. This methodology is the main tool for the qualitative study of the reactive

distillation process which requires minimum initial information. Also it is based on both the

thermodynamic topological behavior of reactive system and on the selection of stable

state limits of highest conversion.

Static Analysis was developed by Serafimov et al [12] and has been sufficiently illustrated

for Pisarenko et al [13] and validated in multiple reactive systems [14-16]. Some

considerations should be made to carry out this analysis: (i) the reaction takes place

under equilibrium conditions and (ii) the reactive distillation column operates to total both

reflux and efficiency. In other words (∞/∞) to conditions are considered. T he main

operation parameters are: the flow ratio of distillated to bottom (P/W) and the volume and

localization of the reaction zone.


To carry out the thermodynamic topological analysis of the reactive system their singular

points are characterized as shown Table 8.4 and the corresponding quaternary residue

map curves is obtained as shown in Figure 8.8.

Table 8.4. Singular Points ** - Acetilation System of 1,3-PD* with 2iP13DO*

Component Type Temperature X1 X2

Azeotrope_H2 Heterogeneous O-iBuAld* 61,35 ºC 0,2079 0,7921

iBuAld* Homogeneous 64,10 ºC 1 ---

Azeotrope_H2 Heterogeneous O-2iP13DO* 92,87 ºC 0,7449 0,2551

H2 Homogeneous O 100,00 ºC 1 ---

2iP13DO* Homogeneous 138,12 ºC 1 ---

13PD* Homogeneous 214,40 ºC 1 --- * iBuAld: iso-Butyraldehyde; 2iP13DO: 2-iso-Propil-1,3-Dioxano; 13PD: 1,3-Propanediol; H2O: Water.

** Singular Points at 1Atm.

Figure 8.8. Residue map curves for the reactive system.

Thus, only one distillation region is obtained with a bunch of residue curves starting from

the azeotrope of minimum boiling point (iso-butyl aldehyde-water) and ending in corner of

1,3-propanediol. By direct separation (formulated distilled) three distillation subregions are

obtained while in the case of indirect separation (bottoms formulated) two distillation


subregiones are founded. Thus, ten different possibilities for the feeding ratio of 2-iso-

propyl-1,3-dioxane to water to the reactive distillation column were analyzed (results no

shown). Then, the feeding ratio that generates the highest products distribution (P/W) at a

total conversion of 2-iso-propyl-1,3-dioxane (see Figure 8.9) corresponds to the 2-iso-

propyl-1,3-dioxane/water molar mixture of 0.3776/0.6224, as is shown in Figure 8.10.

Also, based on both the chemical equilibrium and the residue curve maps, it was

determined that the reactive zone must be localized in the stripping section of the reactive

distillation column. Also, the product obtained in the distillated stream is the

heterogeneous azeotrope iso-butyl aldehyde-water.

Then, in order to verify it was possible to obtain the conditions and the trajectory predicted

by the Static Analysis method, simulations at both ∞/∞ (stages/ reflux ) and finite condition

were carried out. Also, it was found that a self-extractive phenomenon (which can be

understood as a non-lineal variation in the relative volatility of a system with the change of

concentrations in a multicomponent mixture, see Figure 8.11) affects strongly the reactive

distillation column performance. The analysis of the isovolatility curves allowed

determining that a redistribution of the feeding streams lead to a high conversion keeping

the configuration obtained by the Static Analysis method.

Figure 8.9. Direct separation with fed 0.377645/0.622355–2iP13DO/Water


Figure 8.10. P/W ratio, Direct Separation (XF: 0.377645/0.622355-2iP13DO/water)

Figure 8.11. iso-Volatility curve (Water–iso-Butyraldehyde–2-iso-Propil-1,3-Dioxane)

In this case, the water required as reactive to carry out the hydrolysis reaction of 2-iso-

propyl-1,3-dioxane is fed in five different stages to the reactive distillation column. For

instance, in the Scenario 1 the reactive distillation column has 45 stages and the water

stream is fed to the stages: 9, 15, 22, 29, 36, and 43. And the respective mass flows are:

7.56, 11.46, 15.36, 19.27, and 23.17 kg/h.


In this way high conversion levels of 2-iso-propyl-1,3-dioxane are obtained. But, because

of the 1,3-propanediol purity ranges 83.3 and 93.9 wt %, a final distillation process is

required in order to achieve a higher purity of 1,3-propanediol. A summary of the main

simulation results for each scenario is given in Table 8.5.

Table 8.5. Summary of the main simulation results for 1,3-propanediol production from

glycerol

Scenario1

Dilgly 13PD2 Organic2 Aqueous1 IP13DO3 13PD3 13PD4

Temperature K 310 310 283.2 283.1 404.2 457.9 486.8

Mass Flow kg/hr 9115.8 8900.1 920.0 9270.3 453. 8 272.6 252.7

WATER 8536.2 8536.2 14.28 8582.9 0 0.752 0

GLYCE-01 578.9 6.45 0.273 6.18 0.273 0.273 0.273

KPNEUMON 0 27.16 0 0 0 0 0

1,3-P-01 0 273.67 1.633 10.95 1.63 251.1 250.83

ACETI-01 0 62.95 0.387 59.07 1.08 1.08 0.980

ETHAN-01 0 20.15 1.335 17.48 3.74 0.003 0

ISOBU-01 0 0 491.32 557.88 0.449 0.183 0

2IP13DOX 0 0 410.77 35.83 446.60 19.24 0.656

Scenario2


Temperature K 310 310 283.2 283.1 409.1 444.4 487.7

Mass Flow kg/hr 6766.1 6594.4 925.9 6847.5 423.5 247.9 245.0

WATER 6186.5 6186.5 15.51 6227.5 0 2.74 0.007

GLYCE-01 578.9 89.0 5.95 83.04 5.95 5.95 5.95

KPNEUMON 0 20.26 0 0 0 0 0

1,3-P-01 0 250.30 1.70 7.54 1.703 238.4 238.4

ACETI-01 0 57.58 0.473 52.84 1.021 0.747 0.686

ETHAN-01 0 10.37 0.560 8.50 1.001 0 0

ISOBU-01 0 0 515.1 442.4 1.515 1 E-03 0

2IP13DOX 0 0 386.6 25.73 412.3 0.017 0.002


Scenario3


Temperature K 310 310 283.2 283.1 411.4 424.7 487

Mass Flow kg/hr 8531.1 8338.9 872.37 8687.6 427.6 253.6 248.4

WATER 7951.5 7951.5 13.77 7995.5 0 5.15 0.027

GLYCE-01 578.9 53.03 2.277 50.7 2.28 2.28 2.28

KPNEUMON 0 23.73 0 0 0 0 0

1,3-P-01 0 259.07 1.535 10.20 1.53 245.6 245.6

ACETI-01 0 60.30 0.374 56.55 0.7 0.59 0.507

ETHAN-01 0 14.28 0.573 12.37 0 0 0

ISOBU-01 0 0 465.07 527.9 0 0.001 0

2IP13DOX 0 0 388.77 34.32 423.1 0.011 0

The final production of 1,3-propanediol is mainly related to the fermentation yield of both

fermentation stages. Thus, while the decreasing order for the final concentration of 1,3-

propanediol after the second fermentation stage was Scenario 2 > Scenario 1 > Scenario

3 (see Table 8.2) and the decreasing order for the fermentative yield of glycerol to 1,3-

propanediol was Scenario 1 > Scenario 3 > Scenario 2 (see Table 8.3), the decreasing

order for the actual production of 1,3-propanediol was Scenario 1 > Scenario 3 > Scenario

2 (see Table 8.5).

Otherwise, high recovery percentages were achieved for iso-butyl aldehyde, indicating

that low requirements of fresh reactive are required. Higher differences are noticed when

the whole technological scheme is analyzed. The maximum global molar yield from

glycerol to 1,3-propanediol was obtained for the Scenario I, while the minimum was

obtained for the Scenario 2. The relative difference between these two scenarios was

5.21 %, which was close to the relative difference for the fermentation yield obtained for

the same both scenarios, 11.14 %. Thus, it can be stated that the technological

performance of 1,3-propanediol production from raw glycerol depends mostly on the

global conversion of substrate to the main product during the fermentation stage. See

Table 8.6.


Table 8.6. Data representing the behavior of the downstream process

Scenario 1 Scenario 2 Scenario 3

Reactive-Extraction

Extraction efficiency (%) 96.00 96.99 96.06

Distribution coefficient 204.09 203.39 204.91

Loading (Z) 0.177 0.182 0.177

Downstream Process

Global 13PD recovery (%) 91.66 95.26 94.80

IBuAld recovery (%) 98.66 99.48 99.54

Global Process

Global process yield from Glycerol to Lactic acid (%) 0.5244 0.4984 0.5134

8.2 Ethanol production Ethanol can be produced from sugarcane [17], corn starch [17], sugar [18], molasses [19],

cassava [20], wheat [21] or lignocellulosic biomass [22-25]. Glycerol fermentation by

Escherichia coli produces a mixture containing predominantly ethanol, acetate, and

succinate, also low amounts of formiate could be produced [26]. Succinate and acetate

are competitive by-products which could eventually decrease the ethanol yield as was

shown in Figure 6.2. Thus, glycerol can be converted into ethanol and either hydrogen or

formiate. The resulting mixture can be easily purified due to the significant

physicochemical differences among its compounds.

The analysis here performed is based on the results presented by Yazdani and Gonzalez,

about glycerol conversion to ethanol by Escherichia coli SY04 (pZSKLMgldA) [27]. Their

experimental study used two approaches: (i) ethanol and H2 co-production, and (ii)

ethanol and formiate co-production. It was found that the maximum theoretical yield in

both cases was 1 mol of ethanol plus 1 mol of either formiate or hydrogen per each mol of

consumed glycerol. As additional information to perform the simulation, the average

molecular formula for E. coli of CH1.9O0.5N0.2

was used [28].

Three different possibilities for ethanol production from glycerol were considered. The first

and second possibilities use crude glycerol (88 wt %) in a fermentation stage at a dilution


of 10 g/L and 20 g/L, respectively. Meanwhile, the third possibility considered pure

glycerol (98 wt %) at 10 g/L. The flowsheet of these three simulated bioprocess for fuel

ethanol production from glycerol using E. coli is shown in Figure 8.12.

In all cases the flowsheet is the same, but the operational conditions are different.

Obtained glycerol from the purification process (88 wt % or 98 wt %) was cooled at 37 ºC

and diluted (10 g/L or 20 g/L) in fresh water at 37 ºC. Then the glycerol fermentation

process was carried out by E. coli SY04 (pZSKLMgldA) [27] and a mixture of ethanol,

formiate, and cells was produced. Cells were withdrawn by centrifugation and an aqueous

stream of ethanol and formiate was obtained. This stream was distilled and ethanol

concentrated in two distillation columns with 40 and 30 stages respectively. Then, an

ethanol stream between 93 wt % and 94 wt % of purity was obtained (concentration near

to ethanol-water azeotrope 95.6 wt %). Finally, ethanol was dehydrated in a molecular

sieve and fuel ethanol was obtained at 99.5 wt %. The main results of this simulation are

shown in Table 8.7.

1 2 3 4 5 6

Glycerol

Ethanol

Solids

Water

Water waste 1

Water waste 2

AdsorbateDiluted Glycerol

Broth

Distillate2

Distillate1

Figure 8.12. Simplified flowsheet of fuel ethanol production from glycerol at 88 wt % and

98 wt %. 1. Mixed tank, 2. Fermentation tank, 3. Centrifuge, 4. First distillation column, 5.

Second distillation column, 6. Molecular sieves.

In general terms the ethanol production has been described as a process composed of

four main stages named: raw material conditioning, fermentation, separation and

dehydration, and waste treatment. A comparison among the ethanol production from

traditional feedstocks (i.e., sugar cane and corn) versus raw glycerol as raw material was

performed. Figure 8.13 shows the required stages for ethanol production from these three

feedstocks.


Table 8.7. Simulation results for fuel ethanol production from glycerol

STREAM

Diluted glycerol

Broth

Distillate 2

Ethanol

From Crude Glycerol at 10 g/L

Temperature (ºC) 37 37 77.9 77.9

Mass Flow (kg/hr) 57633.551 57624.871 292 273.072

Mass Fraction:

Water 0.99 0.9899 0.06 0.005

Glycerol 0.01 0.0002 0 0

E. coli 0 0.0004 0 0

Ethanol 0 0.0048 0.94 0.995

Formiate 0 0.0046 0 0

From Crude Glycerol at 20 g/L


Mass Flow (kg/hr) 28896.25 28881.1230 313 290.508

Mass Fraction:

Water 0.98 0.9801 0.067 0.005

Glycerol 0.02 0.0000 0 0

E. coli 0 0.0006 0 0

Ethanol 0 0.0103 0.933 0.995

Formiate 0 0.0089 0 0

From Pure Glycerol at 10 g/L


Mass Flow (kg/hr) 57530.195 57538.901 317 293.383

Mass Fraction:

Water 0.99 0.9896 0.07 0.005

Glycerol 0.01 0.0000 0 0

E. coli 0 0.0004 0

Ethanol 0 0.0053 0.93 0.995

Formiate 0 0.0047 0 0


Corn Crude Glycerol

Grinding

Cooking

Liquefaction

Evaporation train

Fermentation

Saccharification

Centrifugation

Drying

Enzyme

Enzyme

CO2

Distillers driedgrains

Ethanol

Evaporation

Ethanol

Whole Stillage

Thin Stillage

DIstillation trainDehydration by adsorption

Neutralization

Washing

Distillation

Fermentation

Centrifugation

DIstillation train

Dehydration by adsorption

Biomass

Sugar cane

Grinding

Clarification

Fermentation

Centrifugation

CO2

Evaporation train

DIstillation train

Co-generation

Combustion gases

Steam

Dehydration by adsorptionEthanol

Concentrated stillage

Whole Stillage

Bagasse

Figure 8.13. Stages for ethanol production from sugar cane, corn, and crude glycerol.

Although simulations were carried out for ethanol production from sugar cane and corn

according to the flowsheets shown in Figure 8.14, the most relevant results were obtained

from the economic assessments and then they are discussed in the Section 8.8.2.

On the other hand, based on the possibility of transforming both raw glycerol and biomass

to ethanol, a sustainable production of biodiesel from oil palm was here proposed. Thus,

sustainable biodiesel production can be performed using only oil palm as a single

feedstock. Palm oil extraction produces mainly two lignocellulosic residues: empty fruit

bunches (EFB), produced in the highest amount; and palm press fiber (PPF) resulting

from press cake separation. They have an important lignocellulosic content and low

moisture, thus both residues can be used as feedstock for bioethanol production [24].

Besides, glycerol is the main by-product of biodiesel production and it can also be

transformed to ethanol. Thus, both oil extraction residues (EFB and PPF) and raw

glycerol are used as feedstock to produce the ethanol required to carry out the

transesterification reaction with palm oil.


Figure 8.14. Simplified flowsheet for ethanol production from: (A) sugar Cane: 1.

Washing tank. 2. Mill. 3. Clarifier. 4. Rotary Filter. 5. Fermentator. 6. Centrifuge. 7.

Absorption column. 8. Concentration column. 9. Rectifying column. 10. Molecular sieves.

11. Evaporator. 12. Boiler. 13. Turbo-generator. (B) Corn: 1. Washing tank. 2. Mill. 3.

Liquefaction reactor. 4. SSF Reactor. 5. Absorption column. 6. Concentration column. 7.

Rectification column. 8. Molecular sieves. 9. First evaporation train. 10. Centrifuge. 11.

Second evaporation train. 12. Air dryer.


The first step is the overall extraction process of palm oil from fresh fruit brunches (FFB).

The whole extraction process includes: Pretreatment stage, where FFB are first cooked

using saturated steam in order to prepare fruit to a subsequent remotion from brunches,

and then they are digested in a cylindrical vertical tank at 100°C, obtaining a separation of

pulps from nuts. Extraction stage, mashed fruits are passed through a screw pressing,

where crude oil is splited up from cake. Refining stage, is made, first decanting crude with

hot water at 90°C, obtaining a decanter cake from additional ,oil is recover. Clarified oil

contains 1% of water, for that reason, must be dried under vacuum conditions before be

stored in oil tanks. Also, obtained press cake is treated, in order to obtain Palm Press

Fiber (PPF) and nuts, these vegetable wastes are used to extract palm kernel oil (PKO)

and palm kernel cake (PKC).

Then, EFB and PPF obtained in the oil extraction process is used in bioethanol production

from lignocellulosic biomass, composed up to 75% of cellulose and hemicellulose. The

overall process usually includes five main steps: biomass pretreatment, cellulose

hydrolysis, fermentation of hexoses, separation and effluent treatment. In first step,

feedstock is pretreated, because composition of this biomass, containing up to 75% of

cellulose and hemicellulose, it should be broken down into fermentable sugars able to be

converted into ethanol and other products [23]. Among available pretreatment methods, in

this work is used a diluted acid pretreatment with sulphuric acid 1-10% at 121 °C [29], in

order to hydrolyze hemicellulose, producing Hexose and Pentose. In a previous work

Cardona et al [18], showed a very promising integrated configuration for bioethanol

production from an energy viewpoint [18, 30] known as simultaneous saccharification and

co-fermentation (SSCF). In this configuration the hydrolysis of cellulose, the fermentation

of glucose released, and the fermentation of pentoses present in the feed stream is

simultaneously accomplished in a same single unit, using a genetically modified

Zymomonas mobilis, Culture broth exiting SSCF bioreactor has an ethanol concentration

of about 6% weight. This stream is concentrated up to 92% in two distillation columns.

The dehydration of ethanol is made by adsorption with molecular sieves. Stillage obtained

from the bottoms of concentration column is evaporated to reduce its volume and

diminishing the costs of its further treatment and the lignin is separated using

centrifugation.


Otherwise, biodiesel is produced from refined palm oil obtained in extraction section. This

vegetable oil is composed by a mixture of triglycerides, being major compounds:

Tripalmitin, Triolein and Trilinolein. They are transeterified, reacting with ethanol using

potassium hydroxide as catalyst. This process is carried out with an integration approach

named multistage reactor-extractor. This process combines the chemical reaction and

liquid–liquid extraction, achieving high selectivity, conversion, productivity, and purity [31].

In this way two main streams are obtained: biodiesel-enriched liquid phase (65% of ethyl

esters), continuously removed from the reactor-extractor and sent to a separation unit

where ethanol is recovered, and glycerol-enriched liquid phase (44% of glycerol) [32].

Finally, an additional amount of ethanol can be produced from raw glycerol as it was

above described.

Simulations are based on Colombian palm industry conditions, reported by Gutierrez et al.

[24] with an average installed processing capacity of crude palm oil of 122 tonnes per day

of FFB. And the lignocellulosic residues are 28.06 tonnes per day of EFB, and 17.98

tonnes per day of PPF. Thus, the total crude palm oil is 21.76 tonnes per day. To carry

out the simulation of integrated biodiesel production process some particularities of each

stage must be considered, which are described as follows: Extraction process considers

the composition data reported by Abdul Aziz et al. [33, 34], and Wan Zahari and Alimon

[35], for both lignocellulosic residues, EFB and PPF, obtained during palm oil extraction,

also the yield extraction was based on the reported data by Prasertsan and Prasertsan

[36], for processing of FFB to crude palm oil.

Ethanol production process from lignocellulosic biomass was analyzed in a previous work

[30], and a brief description of the main processing units is given: pretreatment of

lignocellulosic biomass, enzymatic hydrolysis, and co-fermentation processes were

simulated based on a stoichiometric approach. Thus, lignocellulosic biomass was

converted into glucose and pentoses, which after were converted into cell biomass,

ethanol, and fermentation by-products.

Biodiesel production process from palm oil was previously analyzed by Gutiérrez et al.

[24] in an integrated raw-ethanol production process from lignocellulosic biomass, where

a kinetic approach and a multi-stage reactor–extractor were used. This kinetic model

considers a serial reactive system which transforms triglycerides and ethanol into an ethyl


ester (biodiesel) molecule and either diglycerides, monoglycerides, or glycerol. Figure

8.15 shows jointly the ethanol production process from lignocellulosic biomass, the

biodiesel production process from palm oil, and the ethanol production process from

glycerol.

Here, the process integration is made at different levels to increase efficiency and

productivity. First level is given for two individual processes: i) Integration reaction-

reaction for the process of producing ethanol from biomass feedstock EFB and PPF. In

this process the reaction of hydrolysis of cellulose is carried out simultaneously with the

fermentation of pentoses and hexoses in the SSCF process. ii) Integration reaction-

separation of the process of producing biodiesel from palm oil, which uses a multistage

extractor reactor. Second level of integration takes place between the processes of

biodiesel and ethanol production using lignocellulosic biomass and glycerol. Second level

uses a totally integrated configuration, using oil palm residues EFB and PPF, which are

proposed as raw materials for ethanol production. They enter the pretreatment reactor

where react with dilute acid at high pressure. Then, the pretreated lignocellulosic biomass

undergoes the transformations described in Bioethanol Production from lignocellulosic

biomass Section obtaining dehydrated ethanol with purity greater than 99.5% by weight.

This stream of ethanol, is mixed with an incoming one form (Bioethanol from glycerol)

section, where process crude glycerol is first purified, finally, the crude glycerol is first

refined, and then converted to ethanol by mean of an Escherichia coli strain in a

fermentative process [37]. Final ethanol mixture with a purity of 99.5 % weight, along with

crude oil, is fed to a multi-stage reactor–extractor (Biodiesel from palm oil where

transesterification reaction is continuously accomplished by reactive extraction process

using KOH. Also, main input data and operation conditions used in the simulation process

are shown in Table 8.8.

Thus, sustainable biodiesel production from oil palm was simulated considering jointly

four processes named: palm oil extraction and refining, biodiesel production, and ethanol

production from two feedstocks, lignocellulosic residues and raw glycerol. The

corresponding simulation results for the main process streams are shown in the Table

8.9. Due to extraction process is not showed in the Figure 8.15 the main feed streams are

EFB, PPF, and palm oil; and the other feed streams are service fluids and catalytic


agents. In this sense, the main product streams are biodiesel and ethanol, and also some

waste water streams are obtained.

Figure 8.15. Flowsheet for the integrated process of combined biodiesel and bioethanol

production. Syrup (concentrated sugars in water). (1) Pretreatment reactor, (2) Washing,

(3) Ionic exchange, (4) Simultaneous saccharification and co-fermentation, (5)

Concentration column, (6) Rectification column, (7) Molecular sieves, (8) Evaporation

train, (10) Centrifuge, (11) Multi-stage reactor–extractor, (12) Distillation column for

biodiesel purification, (13) Distillation column for glycerol purification, (14) Neutralization

tank, (15) Centrifuge, (16) First distillation column, (17) Washing tank, (18) Evaporation

column, (19) Second distillation column, (20) Mixed tank, (21) Fermentation tank, (22)

Centrifuge, (23) Third distillation column, (24) Fourth distillation column, (25) Molecular

sieves.


Table 8.8. Main input data and operation conditions used in the simulation process.

Feature Description

Raw materials composition

FFB: Cellulose 18,38; Hemicellulose 12,52; Lignin 9,05; Others 9,05;

Moisture 56,84. EFB: Cellulose 15,47; Hemicellulose 11,73; Lignin 7,14; Ash 0,67;

Moisture 65,00. PPF: Cellulose 24,00; Hemicellulose 14,40; Lignin 12,60; Ash 3,00; Oil

3,48; Others 2,52; Moisture 40,00.

Ethanol production

from lignocellulosi

c biomass

Lignocellulosic biomass pretreatment: H2SO4

Simultaneous saccharification and co-fermentation: T. reesei

cellulases and recombinant Z. mobillis at 30°C by 144 h. Cellulose and

cellobiose conversion are 80% and 100% respectively. Ethanol and

biomass yield are 92% and 2,7% of theoretical.

diluted at 190°C and

12.2 atm by 10 min. Hemicellulosic conversion 75%.

Ethanol distillation: Two distillation columns at 1,77 atm, where the

final ethanol concentration is 92,3 wt %. Ethanol dehydration: Molecular sieves at 116 °C by 10 min at 1,7 atm.

Final ethanol concentration: 99,5 wt %.

Biodiesel production

from palm oil

Multi-stage reactor–extractor: 5 counter-current stages at 60°C with a

residence time of 6 h. Liquid phase equilibrium was considered. Biodiesel purification: Ethanol recovery by distillation from both light

(biodiesel) and heavy (glycerol) phases.

Ethanol production from crude

glycerol

Glycerol fermentation: E. coli SY04 (pZSKLMgldA) is used in a

fermentation broth of 20 g/L of glycerol (previously purified until 88 wt %)

at 37 °C and pH 7. Ethanol distillation: Two distillation columns at atmospheric pressure,

where the final ethanol concentration is 94,1 wt %.

Ethanol dehydration: Molecular sieves at 116 °C by 10 min at 1,7 atm.

Final ethanol concentration: 99,5 wt %.


Table 8.9. Main process streams for ethanol production from lignocellulosic biomass

Stream

Lignocell Biomass

Broth

Recycled

water for washing

Rectific. Column Distillate

BioEtOH From

Biomass

Biodiesel

Refined Glycerol

BioEtOH From

Glycerol

T (°C) 20 30 77,4 93,4 25 73,7 244,9 77,9

P (bar) 1 1 1,793 1,793 1000 0,2 0,3

Mass flow (Kg/h) 1913,1 4185 667,3 310,8 230,7 947,8 95,8 47,11

Cellulose (%) 18,38 1,52 0,04 -- -- -- --

Hemicellulose (%) 12,52 1,31 0,03 -- -- -- --

Lignin (%) 9,05 3,72 0,09 -- -- -- --

Glucose (%) -- 0,56 0,5 -- -- -- --

Xylose (%) -- 0,67 0,89 -- -- -- --

Water (%) 56,84 80,47 97,78 8,5 0,5 0,002 10,5 0,5

Triolein (%) -- -- -- -- -- 0,6 -- --

Diolein (%) -- -- -- -- -- 0,2 1,2 --

Monoolein (%) -- -- -- -- -- 0,02 0,298 --

Ethanol (%) -- 5,59 0,01 91,5 99,5 1,32 0,002 99,5

Ethyloleate (%) -- -- 0,01 -- -- 98,9 -- --

Glycerol (%) -- -- 0,02 -- -- 0,02 88 --

This configuration considers two simultaneous processes: (i) saccharification and

fermentation, where the cellulose hydrolysis produces glucose, which is assimilated by

the microorganisms and converted into ethanol, and thus the inhibitory effect of glucose

over cellulases are reduced. And (ii) reactive extraction where reaction and separation are

integrated in only one processing unit and then conversion, yield and productivity are

improved related to conventional processes because of the continuous products

extraction. In this way, a high product concentration is obtained in both cases.

A high ethanol/palm oil ratio is used inside the multistage reactor-extractor since the

excess of ethanol leads to a better conversion of feedstock during biodiesel production,

as it was showed by Gutiérrez et al [24], because of the ethanol fed to the reactor-

extractor is a mixture of the ethanol obtained from both production processes

(lignocellulosic biomass and glycerol) and the recycled ethanol during biodiesel


purification. This mixture ensures a high ethanol flow inside the multi-stage reactor–

extractor. Thus, a 99,9 % of triolein conversion is reached with a final ethyloleate purity of

98,4 wt %.

Biodiesel production from oil palm and ethanol production from lignocellulosic biomass

have been individually described and analyzed in many times. An important integration

approaches was recently performed by Gutiérrez et al [24], they considered different

process configurations for heat and mass integration, and results were discussed based

only on data obtained from process simulation. But the integration of these two processes

leaves an unsolved problem, which is the production of low cost glycerol; since its sales

do not represent a significant income for the integrated biodiesel production. In this way,

the integration of a biorefinery that uses crude glycerol as feedstock to produce more raw-

ethanol was analyzed. Thus, it is possible to have the oil palm as single raw material to

produce biodiesel. Glycerol conversion to ethanol was 99,8 % and the yield was 99 % of

theoretical.

8.3 PHB production Glycerol purification, glycerol fermentation (cell growth and PHB accumulation), mass cell

pretreatment, PHB isolation, and PHB purification are the five stages needed for the

process of PHB production from raw glycerol. The purification process of raw glycerol was

described in the Chapter 4. This process includes a methanol recovering which decreases

the purification costs in 37.5% [38].C. necator JMP 134 can synthesize PHB up to 70 wt

% of the cell dry mass from various carbon substrates [39]. The fermentation process is

carried out in two stages, in the first stage cell growth occurs and in the second stage

PHB is synthesized. Air and pure oxygen are fed at the first fermentation tank where the

fermentation broth is saturated between 15 or 20 DOC %, thus PHB accumulation takes

place inside the cell mass [40]. The detailed conditions for the fermentation process are

shown in Table 8.10

After fermentation, the next step is PHB isolation and purification. PHB must be extracted

from the cell cytoplasm. Cell membrane is broken and PHB is dissolved and separated

from the residual biomass. The separation step can be divided in three parts:

pretreatment, extraction, and purification. In the pretreatment step cell disruption is


carried out easily and some alternatives for this step are: heat, alkaline or salt

pretreatment and freezing [41].

Table 8.10. Process conditions for glycerol fermentation.

Feature Description

Feedstock

Glycerol 88 wt% Glycerol 98 wt%

Feed concentration: 170.8 g/L 249 g/L

Feed flow rate: 1000 kg/h 1000 kg/h

Pumped

Outlet Pressure: 25 atm 25 atm

Net Work required: 8.1 KJ/Kg 7.9 KJ/Kg

Sterilization

Temperature: 139 ºC

Heat duty: 443.96 KJ/Kg 429.31 KJ/Kg

Heat exchange

Heat duty: 415.96 KJ/Kg 402.29 KJ/Kg

Required area: 63.62 m 51.34 m2

Cell mass growth

2

Temperature: 35 ºC

pH: 7

Residence time: 21 h.

Aeration: 0.6 vol/(vol*min)

Cell mass concentration: 4.91 wt % or 50.4g/L 4.24 wt % or 44.4 g/L

PHB concentration: 0.44 wt % or 4.5 g/L 0.70 wt % or 7.3 g/L

PHB accumulation

Temperature: 35 ºC

pH: 7

Residence time: 22.5 h.

Cell mass concentration: 7.1 wt % or 73.4 g/L 8.7 wt % or 91.5 g/L

PHB concentration: 2.7 wt % or 27.8 g/L 5.5 wt % or 57.1 g/L

Flow rate: 929.05 kg/h 893.173 kg/h

Some of the different extraction methods to separate PHBs from the cell residual

material are: solvent extraction, digestion, mechanical cell disruption, supercritical fluids

extraction, cell fragility and spontaneous liberation. In Table 8.11 advantages and

disadvantages of the most commonly used PHB extraction methods are listed. Solvent

extraction modifies the cell membrane permeability and the PHB is then dissolved [41].


Some used solvents are: chlorinated hydrocarbon (e.g. chloroform), cyclic carbonates

(e.g., propylene and ethylene carbonates), halogenated solvents (e.g., chloroethanes and

chloropropanes), non-halogenated solvents (e.g., chain (4–10 carbons) alcohols, esters,

amides, and ketones (both cyclic and acyclic compounds)). Digestion can be chemically

or enzymatically performed. Chemical digestion uses different chemical agents to destroy

lipids, carbohydrates, proteins and enzymes. According to the chemical agent used, the

chemical digestion could be: digestion by surfactants (e.g. anionic sodium dodecyl sulfate

(SDS) and synthetic palmitoyl carnitine), by sodium hypochlorite, by sodium hypochlorite

and chloroform, surfactant-hypochlorite digestion, surfactant-chelate digestion, and

selective dissolution of non-PHA cell mass by protons.

The enzymatic digestion uses enzymes to degrade the cell membrane. Some varieties of

proteolytic enzymes have high activities on protein dissolutions and slight effects on PHB

degradation. Enzymatic digestion can be complemented by other extraction methods.

Mechanical cell disruption has been widely used to recover intracellular proteins by

different ways [42-43] such as: bead mill disruption, high pressure homogenization,

disruption by ultrasonication, centrifugation, and chemical treatment. Supercritical fluids

have unique physicochemical properties such as high densities and low viscosities that

make them suitable as extraction solvents. Due to its low toxicity and reactivity, moderate

critical temperature and pressure (31°C and 73 atm), availability, low cost, and

nonflammability CO2

is the most used fluid [44]. This extraction method can also be

combined with NaOH or salt (NaCl) pretreatments to get higher disruption levels [41]. The

cell fragility method takes advantage of the cell fragility shown by some bacteria after

large amounts of PHB accumulation. Other extraction methods use air such as: air

classification and dissolved-air flotation. Finally, purification methods involve a hydrogen

peroxide treatment combined with action of enzymes or chelating agents [41].

Besides, based on the available methods for PHB extraction (see Table 8.11) three PHB

production processes from either raw glycerol (88 wt %) or pure glycerol (98 wt %) were

designed.


Table 8.11. PHB Extraction Methods (Adapted from Jacquel et al [25])

Extraction Methods Advantages Disadvantages Results (wt %)

Solvent extraction

Elimination of Endotoxine/high

purity No polymer degradation

Break PHA granules morphology.

Hazards connected with halogenated

solvents. High price/Low recovery

Purity: 99.5%; Recovery> 90%

Digestion by surfactants

Treatment of high cell densities No

polymer degradation

Low purity/Water waste treatment

needed Degradation of the

polymer

Surfactant: High cell density digestion by

SDS)Purity >95%; release rate >90%

Digestion by NaOCl High purity Degradation of the polymer

Purity: 99%; Recovery: 94%

Digestion by NaOCl and chloroform

Low polymer degradation high

purity

High quantity of solvent needed

Purity: >97%; Recovery: 91%

Digestion by NaOCl and surfactants

Limited degradation/low operating cost

-

Surfactant-EDTA disodium salt. Purity: 98%;

Recovery: 86.6%

Digestion by chelate and surfactants

High purity/low environmental

pollution

Large volume of wastewater Low

degradation of the polymer

Purity: 98.7%; Recovery: 93.3%

Selective dissolution of

NPCM (Non PHB cell mas) by

protons

High recovery and high purity low operating costs

- Purity: 98.7 wt% Recovery: 95.4%

Enzymatic digestion Good recovery High cost of

enzymes Purity: 92.6 wt% Recovery: 90%

Bead mill disruption No chemicals used

Require several passes -

High pressure homogenization

No chemicals used

Poor disruption rate for low biomass

levels Low micronization

Yield: 98% Purity: 95%

Supercritical CO2 Low cost, low

toxicity Low recovery Recovery: 89%

Using cell fragility Use of weak - Purity: 99%


extracting conditions

Recovery: 96%

Air classification High purity Low recovery Yield: 90% Purity: 97%

Dissolved air flotation

No chemicals used

Require several consecutive flotation

steps Purity: 86%

Spontaneous liberation

No extracting chemicals needed

Low recovery (80% cells secretes PHB

granules spontaneously

Yield: 80%

The flowsheets for PHB production are shown in Figure 8.16. The fermentation process

begins with a sterilization of diluted glycerol. Pure glycerol was diluted at 249 g/L and raw

glycerol at 170.8 g/L based on the total glycerol consumption by C. necator [40]. Glycerol

conversion takes place in two continuous fermentation stages for cell growth and PHB

accumulation, with operation times of 21 and 22.5 h, respectively. Total glycerol

consumption is considered on the second fermentation stage. The sterilization and

fermentation stages are common for the three PHB production processes. Table 8.12

shows the first downstream process (see Figure 8.16) and it is based on a variation of the

BIOPOL flowsheet [45-46]. The first step is a thermal treatment at 85 ºC, and then a

digestion process using the pancreatin enzyme Burkholdeira sp. PTU9 and NaOCl is

carried out [47]. This pretreatment causes an appropriate cell disruption releasing the

PHB to the fermentation broth. The digestion product containing between 7 to 9 wt % of

biomass is filtrated and the residual cell mass is withdrawn. The mixture containing the re-

suspended PHB at 5.5 – 5.7 wt % is treated with a hydrogen peroxide solution. Then,

using a flash process the majority of the water content is retired. Finally, PHB at 99.9 wt

% is obtained by spray drying.

Figure 8.16 shows the second downstream process and the process conditions are given

in Table 8.13. After passing through the high pressure homogeniser, the depressurized

stream is centrifuged and the solid product is heated and mixed with diethyl succinate

(DES) in a 1/20 ratio of biomass/solvent. The solvent extraction process takes place by

modification of the cell membrane permeability and PHB dissolution [41]. Residual cell

mass is withdrawn by centrifugation and a mixture of PHB-water is gelled by cooling and

the DES is recovered. Finally PHB at 99.9 wt % is obtained by spray drying.


Figure 8.16 also shows the third downstream process for PHB production, which is

described in Table 8.14. This downstream process uses an alkaline pretreatment with a

NaOH solution. Then, a digestion process is carried out using NaOCl and sodium dodecyl

sulfate (SDS) as detergent. The disrupted cells are centrifuged and PHB is washed with

H2O2

. The obtained mixture is subjected to an evaporation process and most of the water

content is discarded. Finally, PHB at 99.9 wt % is obtained by spray drying.

GLYCEROL

AIRO2

GASES

NaOCl

SOLIDS

H2O2

STEAM

WASTEWATER 1

PHB

WASTEWATER 2

ENZYME

1 23

4 5

6789101112

SOLIDS 2 STEAMPHB

WASTEWATER 2 9

101114DES

STEAM

STEAM

1213

DES

8 7

WASTEWATER 1

6

NaOCl

SOLIDS

H2O2

STEAM

PHB

WASTEWATER 2

SDS6789

101112 NaOH

Ferm

enta

tion

stag

eD

owns

tream

P

roce

ss I

Dow

nstre

am

Pro

cess

IID

owns

tream

P

roce

ss II

I

Ferm

enta

tion

brot

h

(88 wt % or98 wt %)

|

Figure 8.16. Flowsheets for PHB production from glycerol (88 or 98 wt %). Fermentation stage: 1. Pump; 2. Sterilizer; 3. Heat exchanger I; 4. Fermenter I; 5. Fermenter II.

Downstream Process I: 6. Heater; 7. Digestor; 8. Centrifuge; 9. Washer tank; 10. Heat

exchanger II; 11. Evaporator; 12. Spray drier. Downstream Process II: 6. Homogenizer;

7. Centrifuge I; 8. Heat exchanger II; 9. Heat exchanger III; 10. Extractor; 11. Centrifuge

II; 12. Heat exchanger IV; 13. Decanter; 14. Spray drier. Downstream Process III: 6.

Alkaline tank; 7. Digester; 8. Centrifuge; 9. Washer tank; 10. Heat exchanger II; 11.

Evaporator; 12. Spray drier.


Table 8.12. Process conditions for PHB recovery: Downstream Process I

Feature Description Feedstock Fermentation broth from glycerol at : 88 wt % 98 wt %

Heat Pretreatment

Temperature: 85 ºC Residence time: 15 min Heat duty: 109.95 KJ/Kg 114 KJ/Kg

Chemical + enzymatic Digestion

Temperature: 50 ºC NaOCl at 30 wt % Ratio NaOCl/cell mass: 1/2 Enzyme: Burkholdeira sp. PTU9 Residence time: 1 h. pH: 9 Enzyme concentration: 2 wt %

Centrifugation

Residence time: 20 min. Retired products: mass cell, mainly.

H2O2 Concentration: 1.2 v/v % -Water washing Heat Exchanging

Heat duty: 2.305 MJ/Kg 2.227 MJ/Kg Required area: 3.06 m 2.95 m2

Water evaporation

2 PHB purity: 37.7 wt % 53.0 wt %

Spray Drying

Heat duty: 2.1542 MJ/Kg 1.02 MJ/Kg Product Purity: 99.9 wt % 99.9 wt % Flow rate: 24.98 kg/h 48.25 kg/h

Table 8.13. Process conditions for PHB recovery: Downstream Process II

Feature Description Feedstock Fermentation broth from glycerol at : 88 wt % 98 wt % Pumped

Pressure outlet: 70 Mpa Net work required: 239.73 KJ/Kg 249.68 KJ/Kg

High pressure homogenizer

Pressure: 70 Mpa Temperature: 110 ºC Residence time: 45 min Heat Duty: 218.40 KJ/Kg 250.02 KJ/Kg

Centrifugation 1

Residence time: 20 min. Recovered products: solids at 62.5 wt % 65 wt %

Heat Exchanging 2


Heat Exchanging 3

2

Solvent to heat: Diethyl-succinate Heat duty: 0.158 MJ/Kg 0.158 MJ/Kg Required area: 0.5 m 0.43 m2 2


Solvent extraction

Temperature: 110 ºC Pressure : 1 atm Mass ratio of PHB/solvent: 1/20

Centrifugation 1

Residence time: 20 min. Extracted products: cell mass

Heat Exchanging 4

Heat duty: 0.200 MJ/Kg 0.216 MJ/Kg Required area: 19.5 m 18 m2

Decantation

2

Temperature: 25 ºC PHB purity: 38 wt % 41.7 wt %

Spray Drying

Heat duty: 1.13 MJ/Kg 1.09 MJ/Kg Product Purity: 99.9 wt % 99.9 wt % Flow rate: 25.36 kg/h 48.74 kg/h

Table 8.14. Process conditions for PHB recovery: Process III

Feature Description Feedstock Fermentation broth from glycerol at : 88 wt % 98 wt %

Alkaline pretreatment

Temperature: 35 ºC Concentration: 3 M Ratio: 0,4 (Kg of NaOH)/(Kg of mass cell)

Chemical + surfactant Digestion

Temperature: 55 ºC NaOCl at 30 wt % Ratio NaOCl/cell mass: 1/3 Surfactant: Anionic sodium dodecyl sulfate (SDS) Heat Duty: 109.20 KJ/Kg 125.01 KJ/Kg Residence time: 20 min.

Centrifugation 1

Residence time: 20 min. Retired products: mass cell

H2O2 Concentration: 1.2 v/v % -Water washing Heat Exchanging 4


Water evaporation

2 PHB purity: 25.2 wt % 37.2 wt %

Spray Drying

Heat duty: 2.53 MJ/Kg 2.14 MJ/Kg Product Purity: 99.9 wt % 99.9 wt % Feed flow rate: 25.13 kg/h 48.46 kg/h


PHB production from crude glycerol starts with the glycerol purification process. Glycerol

content in the feedstock is 60.05 wt % and the rest is mainly methanol, which is recovered

at 99.9 wt % of purity using an evaporation process. The process continues with

impurities treatment and water evaporation and then the stream containing 80.5 wt % of

glycerol is distilled. Two different operation conditions were used for the molar distillated

ratio: 0.11 and 0.40, to obtain glycerol at 88 and 98 wt %, respectively.

The glycerol fermentation process can be carried out by two ways, using glycerol at 88 or

98 wt %. Each way requires different glycerol concentrations in the fermentation media,

which are 170.8 and 249 g/L, for 88 and 98 wt % of glycerol respectively. These

differences account for the impurities of glycerol at 88 wt % which affect the metabolic

process of C. necator. The diluted glycerol stream is sterilized at 139 ºC and 25 atm, in

both cases. Then, temperature and pressure are fitted to operation conditions (i.e., 35 ºC

and 1 atm). The first fermentation stage is called Cell mass growth, where air and oxygen

are fed to reach the stress conditions. This process is carried out for 21 h at pH 7; then if

glycerol at 88 wt % is used, 50.4 g/L of cell mass and 4.5 g/L of PHB are obtained. When

glycerol at 98 wt % is used, 44.4 g/L of cell mass and 7.3 g/L of PHB are obtained. In the

second fermentation stage the operation conditions are kept equal to the first fermentation

stage, where the residence time is 22.5 h and PHB accumulation occurs. Thus, 27.8 g/L

(for glycerol at 98 wt %) and 57.1 g/L (for glycerol at 98 wt %) of PHB are obtained in the

fermentation outlet stream. The fermentation broth is mainly a mixture of incorporated

PHB in the cell mass and water. To recover the PHB from this broth, three different

downstream processes were considered, and each one was evaluated with glycerol at 88

and 98 wt % (see Figure 8.17).

In the Downstream Process I a heat pretreatment is carried out, which denatures the

genetic material and proteins, and destabilizes the outer membrane of the bacterial cells.

The cell mass is then disrupted and PHB is released using for 1 hour a combined

digestion involving Burkholdeira sp. PTU9 enzyme (2 wt %) and a sodium hypochlorite

solution (30 wt %) in a 1/2 mass ratio of NaOCl/cell mass. Furthermore, solids are

removed by centrifugation, followed by water washing and a peroxide hydrogen

treatment. The resulting mixture is heated and near 90 % of water is evaporated. When

the fermentation process is carried out with glycerol at 88 or 98 wt %, a stream with 37.7

or 53.0 wt % of PHB is obtained respectively. Finally, this steam is spray dried until 99.9


wt % of PHB is obtained. Energetically wise the most efficient process is the one that

uses glycerol at 98 wt % since the other one requires evaporating higher water quantities

in the spray drying process (see Table 8.12). The total energy consumptions were 4.57

MJ/Kg and 3.36 MJ/Kg when glycerol at 88 or 98 wt % was used respectively.

Figure 8.17. Scheme for the simulation procedure to synthesize PHB from crude glycerol.

Raw Glycerol: 60 wt %

Glycerol Purification Process

Crude Glycerol:

88 t %

Pure Glycerol:

98 t %

Fermentation Process in Two Stages

Fermentation Process in Two Stages

Dow

nstream

Process I

Dow

nstream

Process II

Dow

nstream

Process III

Dow

nstream

Process I

Dow

nstream

Process II

Dow

nstream

Process III

PHB 99.9 wt %


The pretreatment step in the Downstream Process II is carried out in a high pressure

homogenizer at 70 Mpa, and 110 ºC for 45 min and then the water excess is extracted by

centrifugation requiring 218.40 KJ/Kg processed. The products stream and the stream

containing the solvent Diethyl-succinate (DES) are mixed at 110 ºC, with a mass ratio

PHB/solvent of 1/20. Solvent extraction takes place and the disrupted cell mass is

retrieved by centrifugation. The resulting mixture is decanted at 25 °C and the recovered

DES is recicled in the extraction process. When fermentation is performed with glycerol at

88 or 98 wt %, a 38.0 or 41.7 wt % of PHB purity is achieved respectively. Finally, this

PHB stream is spray dried up to 99.9 wt % of purity. Using glycerol at 98 wt % implies a

higher energy consumption (2.79 MJ/Kg) than that for glycerol at 88 wt % (2.77 MJ/Kg).

These energy consumptions were calculated as a sum of the main energy consumer units

from Table 8.13.

The Downstream Process III differs from the Downstream Process I in the pretreatment

and digestion steps. Pretreatment is carried out in alkaline media at 35 °C using a solution

of NaOH (3 M) in a NaOH/cell mass ratio of 0.4. Then, the combined digestion process

takes place at 55 °C for 20 min. This process involves anionic sodium dodecyl sulfate as

surfactant and sodium hypochlorite (30 wt %) with a mass ratio NaOCl/cell mass of 1/3.

After water evaporation, PHB at 25.2 or 37.2 wt % of purity are obtained when the

fermentation process is carried out with glycerol at 88 or 98 wt %, respectively. Then

spray drying is carried out and PHB is purified up to 99.9 wt %. Nevertheless, glycerol at

98 wt % takes lower energy consumption than glycerol at 88 wt % (i.e., 4.38 MJ/Kg and

4.98 MJ/Kg, respectively).

8.4 D-Lactic acid production In order to analyze the production process of D-lactic acid from raw glycerol, three main

stages have been distinguished: (i) glycerol purification, (ii) glycerol fermentation, and (iii)

D-lactic acid recovery and purification.

Although lactic acid bacteria have been used for D-lactic acid production from

carbohydrate rich feedstocks, it has also been reported the use of alternative biocatalysts

which are mainly engineered Escherichia coli strains able to produce D- or L-lactic acid

[48-52]. But only a few papers have been published on the use of glycerol as carbon


source for D-lactic acid production [53-54]. For instance, Hong et al. [54] compared eight

bacterial strains for lactic acid production from glycerol. Thus, the strain named AC-521

and a member of E. coli, showed the best performance for a fed-batch fermentation

process. On the other hand, Mazumdar et al. [53] engineered several E. coli strains by

overexpressing pathways involved in the conversion of glycerol to lactic acid and blocking

those leading to the synthesis of by-products as it was above described. In all cases they

used a minimal medium supplemented with sodium selenite, Na2HPO4, (NH4)2SO4,

NH4

Cl, and 20 (or 40 or 60) g/l of pure (or crude) glycerol.

Conventional purification of lactic acid from the fermentation broth could be performed by

two routes: (i) crystallizing and acidifying the previously clarified and concentrated (32 wt

%) fermented liquor; or (ii) crystallizing, filtering, dissolving, and subsequently acidifying

the previously precipitated calcium lactate. But, these conventional routes generate huge

quantities of calcium sulphate cake which is difficult to dispose of [55]. The main

impurities contained in the clarified fermentation broth are residual substrate, color, and

other organic acids. Thus, in order to recover and purify the lactic acid, and also to

remove these impurities from the fermentation broth, several processes have been

proposed. The most remarkable purification processes are: adsorption, electrodialysis,

reactive extraction, and reactive distillation. Here, a brief description of each one is given

and some of their advantages and disadvantages are also discussed.

Recovery of lactic acid from a fermentation broth by adsorption requires special

characteristics of extractants and solid sorbents such as: high adsorption capacity and

selectivity, regenerability, and in some cases biocompatibility with microorganisms. Many

carboxylic acid fermentations operate effectively at a pH > pKa of the acid product; for

example lactic acid (pKa = 3.86) fermentation is typically produced at pH 5–6 [55]. In this

case, agents sufficiently basic to retain a significant capacity several pH units above the

acid’s pKa are recommended. Different basic extractants and polymeric sorbents have

been investigated for the extraction and sorption of lactic acid, but the uptaking degree

depends mainly upon the agent basicity and capacity [56]. For instance, weak base

polymer adsorbents such as: IRA-35 [57], MWA- 1, and VI-15 [58], which did not show

high final purity of lactic acid from fermentation broth. Better results were reported for

other studies [59] which used adsorbents with a water-insoluble macro-reticular gel, or a


weak basic anionic exchange resin with a tertiary amine, or a pyridine functional group, or

a strongly basic anionic exchange resin with quaternary functional groups.

Ion exchange technique has also been studied for the recovery of lactic acid from a

fermentation broth by several ionic exchanger resins such as: poly(4-vinylpyridine) resin

(PVP) [60], IRA-420 [61], IRA-400 [62-63], and IRA-92 [64]. Using IRA-92 (weakly basic

exchanger) and under optimal conditions of the fermentation broth (pH 6.0), lactic acid

was recovered with a yield, purity, and specific productivity of 0.826, 96.2%, and 1.16 g

L.A./(g-resin day), respectively [64]. Otherwise, when IRA-400 (a strong anionic exchange

resin) was used by different authors in an fluidized bed column, 0.18 g L.A./(g-resin) and

0.126 g L.A./(g-resin) [65] were recovered. Similar results have also been reported for

lactic acid recovering by ionic exchange resins, such as: 0.1 [66], 0.18 [67], and 0.2 [68] g

L.A./(g-resin).

By mean of a simulated moving bed (SMB) chromatography process (a continuous

separation process consisting of a circle of chromatographic columns) with a PVP resin,

high product purity (99.9 wt %) and high purification yield (>93%) were achieved [69]. On

the other hand, the PVP’s adsorption capacity was found to decrease about 14 % each

time after base regeneration [70].

An advantage for adsorption on ion exchange resin is its possibility to couple it with the

fermentation process. But in the same way, adsorption and ion exchange technologies

require: (i) regeneration of the ion exchange resin, (ii) adjustment the pH of the fed stream

in order to increase the sorption efficiency, (iii) large amounts of extra-chemicals, and (iv)

treatment and disposal of large quantities of salts and effluents [55]. Besides, the

decreasing on the adsorption capacity for the ion exchange resins has also been reported

[55]. Additionally, because of the low adsorption selectivity of some ion exchange resins,

a further purification by esterification is necessary.

Electrodialysis has been recognized as a promissory technology for lactic acid recovery

from the fermentation broth since the product can be continuously removed maintaining

constant the pH of the medium [71]. Recovery of lactic acid is performed from lactate salts

in a two steps process, a conventional electrodialysis to concentrate and purify the

product, and a bipolar electrodialysis to convert the of lactate salts into lactic acids [55]. In


situ lactate recovery electrodialysis has been used with free and immobilized cells in order

to reduce the product inhibition. But, although the final amount of product was increased

in the medium containing immobilized cells, problems related to deposition and fouling of

bacteria on the membranes were found [55]. On the other hand, complete technological

schemes have also been suggested around electrodialysis process for lactic acid

purification. Bailly et al. [72-73] proposed a process in which a conventional

electrodialysis is utilized prior to an electrodialysis stage which uses bipolar membranes

to increase the concentration of organic acid salts. The same configuration of tow-stage

electrodialysis was reported by Habova et al. [74] as an efficient technique for lactate ions

recovering. The same configuration of two-stage electrodialysis was reported by Habova

et al [74] as an efficient technique for lactate ions recovering, while Li et al. [75] combined

both conventional electrodialysis and bipolar membrane electrodialysis in one laboratory

scale bioreactor. These processes let to have a good pH control, lead to reduce the

generation of troublesome salts, and seem to be economically and environmentally

attractive. But, exploitation of electrodialysis with bipolar membranes will require two

previous stages, named: micro-filtration and monopolar electrodialysis. Despite the

several studies performed in order to improve the eletrodialysis fermentation method,

commercialization of this process has not been reported [55].

Other widely studied alternative to recover lactic acid from a fermentation broth is the

reactive extraction. This process uses the reaction between extractants which are in the

organic phase and the extracted materials which are in the aqueous phase. Then, the

complexes formed during the reactions are solubilized in the organic phase. The most

used extractans for the reactive extraction of carboxylic acids are hydrocarbon,

phosphorous, and aliphatic amine [76]. Thus, three categories of extractions are

recognized: (i) extraction by solvation with carbon-bonded oxygen-bearing extractants, (ii)

salvation with phosphorous-bonded oxygen-bearing extractants, and (iii) proton transfer

or ion-pairing formation with high molecular weight aliphatic amines and their salts [77].

But, although only the last two extractants have been mainly used in the recovery of

carboxylic acids, the best extractabilities have been noticed for aliphatic amines. This

attribute is due to the behavior of the acid proton during the transfer from an aqueous

phase to an organic solution. That is, meanwhile the measures of extractability in systems

containing oxygen-bearing extractans are given by both the acid strength in the aqueous

solution and the hydrogen bond in the organic solution; in the case of using aliphatic


amines, the extracted compounds are more stable ammonium salts. The reactive

extraction process of carboxylic acids with tertiary amine extractants is composed of three

sequential steps: dissociation of carboxylic acid, proton transfer to the amine, and

recombination of ammonium salt [77]. The following reaction describes the overall

process, but its stoichiometry varies with several factors, such as: the property and

concentration of amine, acid, and diluent.

R3N + HA- ↔ R3

NHA (8.15)

A successful reactive extractive process depends mainly on a high distribution coefficient

for the lactic acid (Kd

), and also the extractant should have both low water solubility and

low distribution coefficient for the impurities. The distribution coefficient is defined, as ratio

of the lactic acid concentration in the solvent phase to lactic acid concentration in the

aqueous phase, as shown in equation (8.16).

phase aqueous in the in theLA ofion Concentratphase organic in the in theLA ofion Concentrat

=dK (8.16)

Other two important parameters used to evaluate the performance of the reactive

extraction process are the extraction efficiency (E) and the loading (Z), as shown in

equations (8.17) and (8.18), respectively.

100LA ofion concentrat Initial

LA) ofion concentrat Reffinate -LA ofion concentrat (Initial(%) ×=E (8.17)

amine ofion concentrat Initalphase organic in theLA ofion Concentrat

=Z (8.18)

Although long chain aliphatic amines are effective extractants of carboxylic acid from

dilute aqueous solutions, these extractants must always be used dissolved in organic

diluents due to their physical properties such as viscosity, density, and corrosivity [78].

Solvents containing functional groups which interact strongly with complex are called as

Active Diluent (e.g., 1-octanol), while the solvents with low interaction level with complex


are called Inert Diluent (e.g. n-hexane). Thus, nature of the organic diluents affects not

only the basicity of amines but also the behavior of the extraction process [78].

Extractants, mainly ternary amines, including tri-n-octylamine (TOA), tripropylamine

(TPA), tributyl amine (TBA), trilauryl amine, tri-n-butyl phosphate (TBP), triisooctylamine,

Alamine 336 have been reported for lactic acid recovery from aqueous medium; besides

inert and active diluents such as: hexane, heptane, xylene, chloroform, chlorobenzene,

chlorobutane, octanol, decanol, dodecanol, oleyl alcohol, tributyl phosphate,

methylisobutylketone, and methylene chloride-n-hexane, have also been used for this

process [55].

Recovery of lactic acid by reactive extraction is performed through the formation of an

acid-amine complex according to equation (8.15). A second step of regeneration is

required in order to reverse this reaction and to recover both the acid product phase and

the extractant phase available to be recycled. Regeneration could be carried out through

backextraction into an aqueous phase [78] by two approaches, swing either temperature

of diluent composition which leads to changes in the equilibrium relationship. Moreover,

recovering of lactic acid from a loaded solvent phase can also be performed using

solutions of NaOH and HCl [79-80]. One of the most suitable techniques for the

regeneration process is the so-called temperature-swing regeneration [78], where the

extracted stream is mixed with a fresh aqueous stream at a higher temperature to

produce an acid-laden aqueous product and an acid-free organic phase.

Lactic acid can also be obtained by a sequential process containing esterification of crude

lactic acid, distillation of ester, and hydrolysis of ester by reactive distillation in order to

obtain the preceding alcohol and lactic acid. Among the different alcohols analyzed for the

esterification process, methanol appears to be the most suitable one because of the

relatively low boiling of both methanol and methyl lactate. This implies lower energy

consumption for heating is the subsequent processes (i.e., distillation of ester and

reactive distillation). A typical configuration analyzed for lactic acid purification through

reactive batch distillation includes two columns reactants separation from the product and

two reboilers for esterification reaction and hydrolysis reaction [81-83]. In most of the

studies performed on lactic acid purification by reactive distillation, the cation exchange

resin Dowex 50W has been used for both esterification reaction and hydrolysis reaction

[81-84]. The main drawbacks for implementing this technology at industrial scale are not


only the low conversion levels on the esterification reaction (around 50%) but also the

high energy requirements to evaporate a mixture of produced ester and the water present

in the fermentation broth.

Joglekar et al. [55] drew four possible routes for lactic acid purification from a fermentation

broth considering both different fermentation modes and downstream processes as

shown in Table 8.15. Also, the purification costs were estimated, for an assumed

production of 1000 Tons of lactic acid (100 wt %) per year, based on reported costs and

on prices of raw materials and utilities for India. According to the authors, purification cost

of Route 2 was not calculated since the data available on expanded bed ion exchange

adsorption technology is not enough for estimating the costs.

Table 8.15. Downstream processes for lactic acid recovery from a fermentation broth

Route

Fermentation mode

Downstream process

Cost (USD$/LA kg)

1

Continuous

Reactive extraction, re-extraction, esterification, and

hydrolysis by reactive distillation 1.59

2

Continuous

Adsorption/desorption using methanol as eluent,

esterification, and hydrolysis by reactive distillation xxx

3

Batch

Addition of lime, precipitation of calcium lactate,

dissolution in methanol, acidification to separate calcium

sulphate, esterification, and hydrolysis by Reactive

distillation. 1.40

4

Batch

Addition of ammonium hydroxide, micro filtration,

monopolar electrodialysis, bipolar electrodialysis,

esterification with reactive distillation and hydrolysis. 1.74

Here, based on the above reviewed literature, a technological scheme for lactic acid

production is proposed, simulated, and economically assessed as follows. Based on the

results reported by Mazumdar et al. [53], five fermentative scenarios were identified to be

analyzed for the D-lactic acid production from raw glycerol. The scenarios differ on: strain,

substrate concentration, substrate purity, fermentation time, and fermentation stages. In

this way, different values of yield, glycerol consumption, and productivity are reported.

Detailed information for the fermentation stage of each scenario is given in Table 8.16.


Table 8.16. Base information for the glycerol fermentation to D-lactic acid

Scenario

Strain

Glycerol Conc. (g/L)

Glycerol purity for

fermentation

Glycerol consumption

(%)

Molar yield to D-LA

Details

1 LA01(pZSKLMgldA) 20 Pure 100 0.820

Main by product:

Ethanol.

Fermentation time: 36 h

2 LA02(pZSglpKglpD) 20 Pure 100 0.812

Main by product:

Succinic acid.

Fermentation time: 36 h

3 LA02Δdld(pZSglpKglpD) 40 Pure 100 0.833 Fermentation time: 72 h

4 LA02Δdld(pZSglpKglpD) 40 Crude 100 0.859 Fermentation time: 72 h

5 LA02Δdld(pZSglpKglpD) 60 Crude 90 0.934

Two fermentation

stages: 48 h and 36 h,

each one.

The downstream process for D-lactic acid recovery and purification from the fermentation

broth is based on a reactive-extractive process because of its good process

characteristics, such as: low toxicity, low cost, low boiling point, extraction yield, and

recovery yield. Tri-n-octylamine and dichloromethane are used as extractant and active

diluent, respectively. A mixture of tri-n-octylamine diluted in dichloromethane at 0.6 M was

considered for this process and the used weight ratio of fermentation broth to organic

media was 2/1. During the reactive extraction process, a complex molecule is formed

according to equation (8.15) (see Figure 8.18). The back-extraction process is carried out

by a combined effect of changing the extractant concentration (Regeneration Swing

Concentration Process) and changing the temperature profile (Regeneration Swing

Temperature Process), which is reached by a distillation process under vacuum

conditions in order to obtain a highly pure D-lactic acid. Also, distillated water is fed in a

molar ratio of 1/1 respect to the formed complex.

The simplified flowsheet for D-lactic acid production is shown in Figure 8.19. The main

differences among the five scenarios are determined by the fermentation stage, which

leads to different: flows, equipment sizes, and operational conditions. The fermentation

product is a mixture containing mainly organic acids and cell mass, where the cell mass is


withdrawn by centrifugation. The clarified broth is mixed with three organic streams,

where two of those streams correspond to recycled tri-n-octylamine and dichloromethane.

The third one is a mixture of fresh both extractant and diluent because of their lost during

the purification process. Then, the reactive extraction process is performed at 20 °C with

a residence time of 1.5 h and with a mixture of 0.6 M of tri-n-octylamine dissolved in

dichloromethane. Thus, the complex is produced and extracted to the organic phase,

which is distillated in order to remove the remaining water. The distillation product is the

heterogeneous azeotrope of dichloromethane-water. This azeotrope is also obtained from

the distillation of the aqueous phase obtained during the reactive extractive process,

which contains a significant amount of dichloromethane. Then, these two streams are

mixed and treated by decantation in order to recover dichloromethane at 99.7 wt %. The

bottom stream obtained after distillation of the organic phase from the reactive distillation

process is also distillated and the remaining dichloromethane is recovered. This stream is

mixed with the decantation product containing dichloromethane at 99.7 wt %, and the final

obtained purity was 99.8 wt %. The new bottom stream, containing mainly the complex, is

mixed with distillated water and then the back-extraction process takes place by

distillation. The distilled product is tri-n-octylamine at 99.1 wt % while the bottom product

is a mixture of D-lactic acid (85.5 wt %) and water. This last stream is finally purified by

vacuum distillation and D-lactic acid at 99.9 wt % is obtained.

The fermentation processes were simulated using a yielding approach where glycerol is

consumed and products including cell mass are formed according to Table 8.17. The

molecular formula used for E. coli strains was CH1.9O0.5N0.2

[85-86].

Figure 8.18. Complex formed during the reactive extraction process of D-lactic acid


Raw Glycerol

Metanol

Solids

Water

OrganicPhase

Aqueous glycerol

Waste Water 1

Waste Water 2

Glycerol(85 wt %)

E-1 R-1 C-1 Dec-1

E-2

DC-1RII-1

Glycerol(98 wt %)

Adsorbate

Solids

Water

Bottoms-6

Diluted Glycerol

Fermentation Broth

Distillated-6

M-1

F-1Cen-1

DC-6

Fresh TOADCE

RE-1

Aqueous Phase

Bottoms-2

Distillated-2DC-2

OrganicPhase

Dec-2Aqueous

Phase

Bottoms-3

Distillated-3

DC-3

M-2

M-3

Bottoms-4

Distillated-4

DC-4Lactic Acid

DC-5

Distillated-5

Figure 8.19. Simplified flowsheet for D-lactic acid production from raw glycerol. E:


fermentator, RE: Reactor extractor.

Table 8.17. Stoichiometry for glycerol fermentation to D-Lactic Acid by Engineered E. coli.

Scenario

Glycerol

Residual Glycerol

Ac Ac

Succ Ac

EtOH

D-Lac Ac

For Ac

CO

2 Biomass

1 -1 0 0.021 0 0.024 0.82 0 0.045 11.5

2 -1 0 0.044 0.008 0 0.812 0 0.044 6.5

3 -1 0 0.048 0.0076 0 0.833 0 0.048 3.4

4 -1 0 0.045 0.008 0 0.859 0.006 0.045 4.5

5 -1 0.1 0.0414 0.00855 | 0.8406 0 0.0414 5

D-lactic acid production process starts with the purification of raw glycerol up to the

required purity according to each scenario as it was showed in Table 8.16. Detailed

information about the glycerol purification process was previously reported [87-88]. A

summary of the main simulation results for each scenario is given in Table 8.18. The final

production of D-lactic acid is directly related not only to the fermentation yield but also to


the substrate consumption. In other words, while the decreasing order respect to the yield

was: Scenario 5 > Scenario 4 > Scenario 3 > Scenario 1 > Scenario 2, the decreasing

order respect to the D-Lactic acid production was: Scenario 4 > Scenario 5 > Scenario 3 >

Scenario 1 > Scenario 2. This change in the order between the Scenarios 5 and 4 occurs

due to the incomplete consumption of glycerol during the fermentation in the Scenario 5.

Table 8.18. Summary of the main simulation results for D-lactic acid production process

Scenario 1

Dilut gly Lactac1 Lacacid4 Lacacid6 Destwater Lactac9 Lactacp

Mass Flow (kg/hr) 28575.1 28560.0 13959.4 3073.5 83.96 487.3 419.5 Water 27995.5 27995.5 68.2 0 83.96 67.5 0.165 Glycerol 578.9 0 0 0 0 0 0 Formic Acid 0 0 0 0 0 0 0 Ethanol 0 6.95 0.855 0.001 0 0.001 0 Acetic Acid 0 7.9 0.612 0.612 0 0.417 0.004 Succ Acid 0 0 0 0 0 0 0 Lact Acid 0 464.3 0.273 0.273 0 419.3 419.3 Ecoli 0 72.2 0 0 0 0 0 Tri-n-amine 0 0 986.1 986.0 0 0 0 Dichlmethan 0 0 10813.9 1.394 0 0.125 0 Complex 0 0 2085.2 2085.2 0 0 0

Scenario 2


Mass Flow (kg/hr) 28575.1 28531.5 13954.7 3068.9 84.6 484.1 416.3 Water 27995.5 27995.5 68.0 0 84.633 67.97 0.172 Glycerol 578.9 0 0 0 0 0 0 Formic Acid 0 0 0 0 0 0 0 Ethanol 0 16.6 1.28 1.28 0 0.87 0.87 Acetic Acid 0 5.9 0.08 0.08 0 0.079 0.079 Succ Acid 0 459.8 0.27 0.27 0 415.1 415.1 Lact Acid 0 40.8 0 0 0 0 0 Ecoli 0 0 1002.4 1002.4 0 0 0 Tri-n-amine 0 0 10813.5 0 0 0 0 Dichlmethan 0 0 2064.8 2064.8 0 0 0

Scenario 3



Mass Flow (kg/hr) 14359.1 14310.3 14231.1 3081.2 86.89 498.6 421.58 Water 13779.5 13779.5 69.92 0 86.89 70.18 0 Glycerol 578.9 0 0 0 0 0 0 Formic Acid 0 0 0 0 0 0 0 Ethanol 0 18.1 2.702 2.702 0 1.85 0 Acetic Acid 0 5.64 0.161 0.161 0 0.157 0.157 Succ Acid 0 471.7 0.586 0.586 0 426.4 421.4 Lact Acid 0 21.35 0 0 0 0 0 Ecoli 0 0 959.5 959.5 0 0 0 Tri-n-amine 0 0 11073.0 0 0 0 0 Dichlmethan 0 0 2118.2 2118.2 0 0 0

Scenario 4


Mass Flow (kg/hr) 14402.8 14375.6 14243.9 3095.2 89.87 516.36 441.7 Water 13821.2 13821.2 70.13 0 89.87 73.09 0.188 Glycerol 580.6 0 0 0 0 0 0 Formic Acid 0 1.741 0.03 0 0 0 0 Ethanol 0 17.04 2.535 2.535 0 1.759 0.019 Acetic Acid 0 5.956 0.17 0.17 0 0.166 0.166 Succ Acid 0 487.8 0.605 0.605 0 441.3 441.3 Lact Acid 0 28.33 0 0 0 0 0 Ecoli 0 0 901.1 901.1 0 0 0 Tri-n-amine 0 0 11071.9 0 0 0 0 Dichlmethan 0 0 2190.8 2190.8 0 0 0

Scenario 5


Mass Flow (kg/hr) 9622.8 9640.9 14320.4 3087.3 88.01 505.9 432.5 Water 9041.2 9041.2 70.94 0 88.01 71.28 0.114 Glycerol 580.6 56.32 0.066 0.066 0 0.066 0.066 Formic Acid 0 0 0 0 0 0 0 Ethanol 0 15.67 3.3 3.3 0 2.274 0.012 Acetic Acid 0 6.365 0.274 0.274 0 0.267 0.267 Succ Acid 0 477.4 0.902 0.902 0 432.0 432.0 Lact Acid 0 31.48 0 0 0 0 0 Ecoli 0 0 938.8 938.8 0 0 0 Tri-n-amine 0 0 11155.0 0 0 0 0 Dichlmethan 0 0 2143.9 2143.9 0 0 0


Otherwise, for the reactive extraction process no high differences were noticed in terms of

Distribution coefficient or Loading as is shown in Table 8.19. Global recovery of D-lactic

acid was around 90 % respect to its production during the fermentation process. Also,

high recovery percentages were achieved for both tri-n-octylamine and dichloromethane,

indicating that low requirements of fresh both extractant and diluent are required. Higher

differences are noticed when the whole technological scheme is analyzed. The maximum

global molar yield from glycerol to D-lactic acid was obtained for the Scenario 4, while the

minimum was obtained for the Scenario 2. The relative difference between these two

scenarios was 5.93 %, which was close to the relative difference for the fermentation yield

obtained for the same both scenarios, 5.47 %. Thus, it can be stated that the

technological performance of D-lactic acid production from raw glycerol depends mostly

on the global conversion of substrate to the main product during the fermentation stage.

Table 8.19. Data representing the behavior of the downstream process for D-lactic acid

production

Scenario 1 Scenario 2 Scenario 3 Scenario 4 Scenario 5

Reaction-Extraction Process Extraction efficiency (%) 92 92 92 92 92

Distribution coefficient 11.59 11.58 11.68 11.68 11.78

Loading (Z) 0.63 0.62 0.64 0.66 0.65

Downstream Process Global lactic acid recovery (%) 90.32 90.3 89.36 90.48 90.51

Tri-n-octylamine recovery (%) 99.998 99.992 99.998 99.998 99.998

Dichloromethane recovery (%) 99.98 99.57 99.99 99.99 99.93

Global Process Global process yield from

Glycerol to Lactic acid (%)

71.39

70.68

71.75

75.14

73.55


8.5 Succinic acid production In general terms, the downstream process for succinic acid purification and recovery from

the fermentation broth starts with a cell separation by centrifugation or microfiltration, and

in some cases an additional ultrafiltration process is used in order to separate residual cell

mass, proteins, and other fermentation supernatants. For succinic acid recovery and

purification different alternatives have been proposed, such as: precipitation with

ammonia or calcium hydroxide, electrodialysis, reactive extraction, and sorption/ion

exchange. Here, the advantages and disadvantages of these alternatives are briefly

discussed.

Industrially the most used method for recovery of carboxylic acids from a fermentation

broth is the precipitation process with calcium hydroxide or calcium oxide, especially for

the cases of lactic and citric acids. Addition of calcium hydroxide or calcium oxide to the

clarified fermentation broth leads to calcium salt formation of succinic acid, which are

filtered off and treated with concentrated sulfuric acid. This last addition generates

calcium sulfate (CaSO4

), gypsum, in an equimolar amount. Then, free succinic acid is

purified by active carbon or ion exchange, and finally the product is further concentrated

and crystallized by evaporation. From a commercial view of point, this purification way

cannot be used because high amount of calcium sulfate are by-produced as a waste a

adequate disposal is required [89-90]. Additionally, the precipitation process requires high

consumption of calcium hydroxide, calcium oxide, and sulfuric acid which cannot be

regenerated or recycled causing high process costs. That is why precipitation with

calcium hydroxide or calcium oxide has been reported as unlikely process for large-scale

production of bio-succinic acid [91]. Precipitation with ammonia has also been reported for

succinic acid recovery at laboratory scale [92-93]. During this process the produced

diammonium succinate must be later treated with sulfate ions, or ammonium bisulfate, or

sulfuric acid at a low pH in order produce both succinic acid precipitate and ammonium

sulfate. Finally succinic acid is obtained after the dissolution in methanol and re-

crystallisation processes while ammonium sulfate can be cracked thermally in order to

produce ammonia and ammonium bisulfate. Although precipitation with ammonia reduces

the amount of waste production, low selectivities have been reported for this purification

alternative [93].


Other alternative for succinic acid purification is the electrodialysis process. Membranes

are charged either with positive or negative groups and selectively allow to cations or

anions passing through the membranes, thus succinate anions are able to passage

through positively charged membranes while sodium cations are repelled. Glassner et al.

[94] reported a total purification yield of 60% for a desalting electrodialysis combined with

a water-splitting electrodialysis. This process requires both a set of chelating ion-

exchange columns to replace the divalent cations of the succinate salt with sodium ions

and a bipolar membrane water-splitting dialysis to obtain succinic acid from succinate.

Then, after electrodialysis and ion exchange process, evaporation of water and

crystallization of the succinic acid are required [94-96]. Electrodialysis is known as an

expensive alternative not only by the high energy consumption but also by the material’s

cost. Besides to the low yield, some other problems such us: low selectivity [97], handle

on binary ions [95], and fouling [98] have also been reported.

On the other hand, since liquid-liquid extraction has shown low distribution coefficients for

carboxylic acids recovery from the fermentation broth [99-100], reactive extraction

appears as a better option to increase yield and selectivity to organic acids from an

aqueous phase [101]. Mixtures of amines (reactive components) dissolved in non-water

miscible organic solvents have been widely studied for carboxylic acids recovery [102-

103]. Amine reacts with the succinic acid thought out a proton transfer or ion pair

formation mechanism depending on the type of amine and the organic solvent [104].

Long-chain aliphatic primary, secondary, and tertiary amines have been proposed for

succinic acid extraction [105-111], but in these cases only the undissociated acid can be

extracted. Otherwise, although quaternary amines can extract the dissociated and the

undissociated succinic acid, its regeneration is difficult by back extraction. Thus, ternary

amines have been the most used for carboxylic acids extraction from an aqueous solution

[112] using solvents such as: octanol, xylene, heptane, kerosene, methylenchloride or

nitrobenzene [105-108, 111, 113]. Mixtures of amines such us: tripropylamine/

trioctylamine or trialkylamines have also been reported [108, 114]. On the hand, different

operational alternatives for the reactive extraction process of succinic acid with amines

have been proposed. For instance, Huh et al. [115] studied the removal of by-products

and impurities present in the fermentation broth such as: acetic acid, pyruvic acid, and

salts, by mean a pre-treatment step of reactive extraction with trioctylamine. Then the

aqueous succinic acid is purified up to 99.8 wt % by an evaporative crystallizer with a


purification yield of 71.3 %. Finally, Kurzrock and Weuster-Botz [101] stated that if the

recycling of the costly amines is done efficiently, it is very likely that optimized reactive

extraction processes may be applied in the near future for industrial production of bio-

succinic acid.

In general terms, the final step of the succinic acid purification process is an ionic exchanger unit in which the residual cations and anions are removed. Different kinds of

exchange resins have been investigated in order to produce succinic acid at high purity

from either the fermentation broth or succinate. Some examples are the alkaline

anionexchange resin (NERCB 04) [116] and the H-type strongly acidic cation-exchange

resins [117]. On the other hand, mesoporouses silicas (SBA-15) functionalized with

primary, secondary and tertiary amino-functional silanes were reported to be able for for

the isolation of pyruvic and succinic acid from fermentation broth [118]. Ion-exchange

must be only considered as an additional purification step for succinic acid recovery

because of its low both selectivity and yield [119].

In this case, three different types of strains are here analyzed for the succinic acid

production from raw glycerol. The first one is the Escherichia coli recently reported by

Blankschien et al. [120], the second strain is the Mannheimia sp. Pasteurellaceae

reported by Scholten and Dagele [121], and the last one is the Anaerobiospirillum

succiniciproducens reported by Lee et al. [122]. In the case of Mannheimia sp.

Pasteurellaceae three scenarios are studied while for Escherichia coli and

Anaerobiospirillum succiniciproducens one scenario is analyzed in each case, as shown

in Table 8.20. Although it can be observed that three different qualities of glycerol are

considered for the fermentation stage (i.e., 76, 90, and 98 wt %), a unique raw glycerol is

considered as feedstock. A typical composition of a raw glycerol stream obtained from a

biodiesel production process is: methanol 32.59 wt %, glycerol 60.05 wt %, NaOCH3

2.62

wt %, fats 1.94 wt %, and ash 2.8 wt % [123]. This stream must be purified up to the

specified concentration established in Table 1 for the fermentation stage. The purification

process previously studied in the Chapter 4. As result of the purification process, a

glycerol stream at 80 wt % is obtained. Thus, the glycerol concentration required for the

Scenario 4 is obtained. Finally, in order to purify the glycerol stream up to 90 or 98 wt %,

a distillation process is carried out.


Table 8.20. Base information for the glycerol fermentation to succinic acid

Scenario

Strain

Glycerol Conc. (g/L)

Glycerol purity for

fermentation

Glycerol consumption

(%)

Molar yield to

S.A. 1 Escherichia coli 20 98 wt % 96 0.522 2

Mannheimia sp. Pasteurellaceae

9.6

98 wt %

55.2

0.287

3


8.3

90 wt %

75.9

0.542

4


9.1

80 wt %

71.4

0.453

5

Anaerobiospirillum succiniciproducens

6.5

98 wt %

58.5

0.344

Then, the fermentation process can be performed, but due to these scenarios differ on:

strain, substrate concentration, substrate purity, and fermentation time, different

composition profiles are obtained in the fermentation broth according to the fermentation

stoichiometry reported for each case and shown in Table 8.21. The fermentation broth is

then clarified by a centrifugation process where the cell mass is withdrawn.

Table 8.21. Stoichiometry for glycerol fermentation to succinic acid by Engineered E. coli.

Scenario Purity of Gly Glycerol

Res Gly

Succinic acid

Acetic acid

Lactic acid

Formic acid Biomass

(wt %) 92.09 92.09 118.09 60.05 90.08 46.03 1 99 -1 0.040 0.544 0.064 0.000 x 0.066

2 99 -1 0.448 0.520 0.048 0.000 0.083 0.044 3 90 -1 0.241 0.714 0.055 0.000 0.072 0.044 4 76 -1 0.286 0.634 0.051 0.011 0.066 0.044 5 99 -1 0.415 0.588 0.020 x x 0.032

Recovery and purification of succinic acid from the clarified fermentation broth is based on

the downstram process recently reported by Huh et al. [124]. During the so-called

complex process, the by-produced acids are removed selectively from the fermentation

broth by reactive extraction, followed by a vacuum distillation step where the volatile

impurities are removed. Then a crystallization process is performed and succinic acid is

concentrated in the order of five to six folds. Finally, another crystallization process is

carrying out at pH 2.0 and 4 °C in order to obtain succinic acid highly purified.


The pretreatment step of the reactive extraction process selectively removes

contaminated organic acids from the dilute fermentation broth using tri-n-octylamine

(TOA) and 1-octanol. Since TOA only extracts the undissociated form of carboxylic acids

[110, 124], the selective removal of specific acids from the fermentation broth is made

possible by using different degrees of dissociation of each acid with the pH. Although the

distribution coefficient values decreased with the increase of pH for succinic acid, pyruvic

acid, and acetic acid, the distribution coefficient of succinic acid at pH values between 4.0

and 5.0 is close to 0. Thus, the increase of the dissociated acid concentration leads to the

reduction of succinic acid extraction. In this, case a multi-stage reactive extractive process

was analyzed, but no significant improvements were noticed after each stage. Besides,

the author recommended a one stage reactive extraction process in order to do economic

the removal of succinic acid from the fermentation broth. Then, by mean a vacuum

distillation process volatile impurities such as acetic acid and formic acid are effectively

removed, and thus the pretreated fermentation broth is concentrated five- to six-fold by a

crystallization process at low temperature (4 °C) and in an acid media (pH 2.0). Table

8.22 shows the calculated removal efficiency of carboxylic acids from the fermentation

broth based on the data reported by Huh et al. [124]. Thus, applying this complex process

all by-produced carboxylic acids are effectively withdrawn with a succinic acid lost of 27%.

Table 8.22. Removal efficiency (%) of the carboxylic acids from the fermentation broth.

Fermentation product Reactive extraction

Vaccum distillation Cristallization

Succinic acid 0.45 0.00 26.58

Maleic acid 87.50 0.00 100.00

Acetic acid 44.44 90.00 100.00

Pyruvic acid 27.14 11.76 99.78

Fumaric acid 87.50 0.00 100.00

On the other hand, due to the fermentation products here considered are: succinic acid,

acetic acid, and formic acid according to Table 8.21, it is required to generate the complex

molecules as the corresponding reaction products of TOA with of carboxylic acids.

Complex I (see Figure 8.20.a) is the reaction product of succinic acid with TOA, while the

Complex II (see Figure 8.20.b) and Complex III (see Figure 8.20.c) are the reaction

products of formic acid and acetic acid respectively.


The simplified flowsheet for succinic acid production is shown in Figure 8.21. As it was

observed for the D-lactic acid production processes, the main differences among the five

scenarios for the succinic acid production processes are determined by the fermentation

stage. Thus different flows, equipment sizes, and operational conditions are obtained for

each scenario. Glycerol is first purified up to the required quality according to Table 8.20

and then the fermentation occurs under the stoichiomentry presented in the Table 8.21.

The fermentation product contains a mixture of cell mass and organic acids such as

succinic acid, acetic acid, lactic acid, and formic acid. The biomass produced during the

fermentation stage is withdrawn by centrifugation and the obtained clarified fermentation

broth is mixed with both the fresh and the recycled streams containing TOA and octanol.

Then, the resulting mixture is subjected to the reactive extraction process.

During the reactive extractive process the complex molecules are produced from each

carboxylic acid according to the Figure 8.20, and then most of the succinic acid complex

is contained in the aqueous phase while the rest of the impurities are extracted to the

organic phase. This organic phase is subjected to a back extraction process by mean a

distillation column in which some water is withdraw and most important the used TOA is

recovered. This stream is again distilled and most of the carboxylic acids are discarded

while a mixture of octanol and TOA containing low quantities of succinic acid and acetic

acid is recycled in order to be used in the reactive extractive process. On the other hand,

the aqueous phase containing most of the succinic acid is first distilled not only water but

also acetic acid and and a few amount of octanol are discarded by distillated. The

bottoms product is then subjected to a crystallization process and succinic acid at high

purity is obtained.

The main simulation results for each scenario are shown in Table 8.23 and significant

differences for the final production of succinic acid can be observed among the five

analyzed scenarios. For instance based on Table 8.21, the final production of succinic

acid depends mainly on the succinic acid yield (increasing order: Scenario 2 < Scenario 1

< Scenario 5 < Scenario 4 < Scenario 3) and also on the glycerol consumption (increasing

order: Scenario 2 < Scenario 5 < Scenario 4 < Scenario 3 < Scenario 1). This higher

dependence on the succinic acid yield can be explained because of the increasing order

of succinic acid production (Scenario 2 < Scenario 1 < Scenario 5 < Scenario 3 <

Scenario 4) was most similar to the succinic acid yield than the glycerol consumption.


Figure 8.20. Reaction complexes of succinic acid, formic acid and acetic acid with TOA.

Raw Glycerol

Metanol

Solids

Water

OrganicPhase

Aqueous glycerol

Waste Water 1

Waste Water 2

Glycerol(85 wt %)

E-1 R-1 C-1 Dec-1

E-2

DC-1RII-1

Glycerol(98 wt %)

Adsorbate

Solids

Water

Bottoms-4

Diluted Glycerol

Fermentation Broth

Distillated-4

M-1

F-1Cen-1

Fresh TOA

Octanol

RE-1

Aqueous Phase

Bottoms-2

Distillated-2DC-2

Bottoms-3

Distillated-3

DC-3

M-2

Succinic acid

Waste Water 3

DC-4

Cry-1

Figure 8.21. Simplified flowsheet for succinic acid production from raw glycerol. E:


fermentator, RE: Reactor extractor, Cry: Crystallizator


The change between the Scenarios 3 and 4 occurs due to the difference in the glycerol

consumption is higher than the succinic acid yield. For the reactive extractive process was

noticed that most of the fermentation by-products such as formic acid and acetic acid

were selectively removed from the clarified broth with extraction efficiency around 88 %

for formic acid and 90 % for acetic acid, while the losses for succinic acid were lower than

1.4 % as shown in Table 8.24. The good performance of the downstream process is

stated because of the high recovery levels not only for succinic acid (> 98 %) but also for

both tri-n-octylamine and 1-octanol.

Table 8.23. Summary of the main simulation results for succinic acid production process

Scenario 1

Dilutgly Succaci1 Complex2 Aqueous2 Extract2 Concent1 Succaci2

Mass Flow kg/hr 28575.1 28447.3 29668.5 27813.1 1855.5 3448.4 123.6 Water 27995.5 27995.5 27992.7 27659.3 333.4 3298.1 0.003 Glyce-01 578.9 23.16 23.15 23.1 0.072 23.1 X Cell mass 0 10.25 X X X X X Succini 0 403.8 403.4 398.2 5.18 398.2 397.6 Formi-01 0 0 0 0 0 0 0 Aceti-01 0 24.16 2.32 0.097 1.74 X 0.811 Tri-N-01 X X 1.99 1.91 0.08 1.91

1-oct-01 X X 1368.8 2.83 1365.9 trace X Complex X X 1.61

1.61 X X

Complex3 X X 149.1 X 149.1 X X Scenario 2

Dilutgly Succaci1 Complex2 Aqueous2 Extract2 Concent1 Succaci2 Mass Flow kg/hr 59382.1 59490.7 62220.1 58525.4 3694.7 7297.5 117.2 Water 58802.5 58802.5 58796.6 58134.5 662.2 6915.4 0.007 Glyce-01 578.9 259.3 259.3 258.6 0.759 258.6 0 Cell mass X 6.83 X X X X X Succini X 386.0 385.6 380.9 4.7 380.9 379.7 Formi-01 X 24.01 3.00 2.89 0.115 0.033 0 Aceti-01 X 18.12 1.81 1.74 0.069 1.31 0 Tri-N-01 X X 3.86 3.71 0.147 3.71 0 1-oct-01 X X 2737.5 5.62 2731.9 0 0 Complex X X 1.54 0 1.54 0 0 Complex2 X X 181.5 0 181.5 0 0 Complex3 X X 111.8 0 111.8 0 0


Scenario 3 Dilutgly Succaci1 Complex2 Aqueous2 Extract2 Concent1 Succaci2 Mass Flow kg/hr 68527.7 68659.9 61970.6 58546.1 3424.4 7237.9 141.3 Water 67945.3 67945.3 58867.9 58248.9 619.1 6948.8 0.007 Glyce-01 578.9 139.5 120.9 120.5 0.33 120.5 0 Cell mass X 6.83 X X X X X Succini X 530.0 458.7 453.5 5.23 453.5 452.9 Formi-01 X 20.83 2.26 2.17 0.081 0.025 0 Aceti-01 X 20.76 1.80 1.73 0.064 1.3 0 Tri-N-01 X X 26.92 25.96 0.961 25.96 0 1-oct-01 X X 2555.0 5.59 2549.4 0 0 Complex X X 1.83 0 1.83 0 0 Complex2 X X 136.4 0 136.4 0 0 Complex3 X X 111.0 0 111.0 0 0

Scenario 4


Mass Flow kg/hr 61264.8 61370.3 64451.9 60269.7 4182.2 7467.8 143.9 Water 60675.0 60675.0 60689.0 59936.9 752.1 7143.8 0.007 Glyce-01 579.3 165.7 165.7 165.1 0.534 165.1 0 Cell mass 0 6.84 X X X X X Succini 0 471.0 470.5 464.2 6.29 464.2 463.5 Formi-01 0 19.10 2.39 2.29 0.100 0.026 0 Aceti-01 0 29.09 2.91 2.79 0.122 2.088 0 Tri-N-01 X X 13.42 12.86 0.562 12.86 0 1-oct-01 X X 3102.52 5.83 3096.7 0 0 Complex X X 1.88 0 1.88 0 0 Complex2 X X 144.5 0 144.5 0 0 Complex3 X X 179.5 0 179.5 0 0

Scenario 5


Mass Flow kg/hr 87662.1 87767.5 87961.0 87355.2 605.8 10757.5 138.6 Water 87082.5 87082.5 87073.8 86965.4 108.4 10375.8 0.01 Glyce-01 578.9 240.2 240.2 240.1 0.077 240.1 0 Cell mass 0 4.97 X X X X X Succini 0 436.5 436.0 435.4 0.603 435.4 434.8 Aceti-01 0 7.55 0.755 0.752 0.003 0.563 0 Tri-N-01 X X 2.43 2.42 0.011 2.42 0 1-oct-01 X X 456.2 7.82 448.4 0 0 Complex X X 1.74 0 1.74 0 0 Complex3 X X 46.58 0 46.58 0 0


Finally, the maximum global yield from glycerol to succinic acid was obtained for the

Scenario 4, while the minimum was obtained for the Scenario 2 as shown in Table 8.24.

This behavior agrees with the order of succinic acid production above described. The

relative difference between these two scenarios was 18.24 %, which indicates that the

fermentation process has a high influence on the final production of succinic acid.

Table 8.24. Data representing the behavior of the downstream process for succinic acid

production


Reaction-Extraction Process

Extraction efficiency of SA (%) 1.28 1.22 1.14 1.34 0.14

Extraction efficiency of FA (%) NA 87.98 87.95 88.02 NA

Extraction efficiency of AA (%) 90.40 90.38 90.36 90.42 90.04

Distribution coefficient for SA 0.168 0.168 0.170 0.168 0.172

Distribution coefficient for FA NA 99.863 107.627 91.304 NA

Distribution coefficient for AA 121.648 128.240 138.222 117.236 1120.455

Loading for SA (Z) 0.118 0.054 0.064 0.063 0.041

Loading for FA (Z) NA 0.618 0.498 0.528 NA

Loading for AA(Z) 0.980 0.367 0.465 0.590 0.915

Downstream Process

Global SA recovery (%) 98.62 98.38 98.76 98.56 99.76


1-octanol recovery (%) 99.66 93.30 95.98 95.99 94.59

Global Process

Global process yield from Glycerol to SA (%)

0.517

0.493

0.589

0.603

0.565

S.A. succinic acid; F.A.: formic acid; A.A.: acetic acid


8.6 Propionic acid production Propionic acid is a naturally occurring carboxylic acid which, in the pure state, is a

colorless, corrosive liquid with an unpleasant odor. Propionic acid is used in the

manufacture of herbicides, chemical intermediates, artificial fruit flavors, pharmaceuticals,

cellulose acetate propionate, and preservatives for food, animal feed, and grain.

As it was performed for D-lactic acid and succinic acid production, the whole technological

scheme for propionic acid production was divided in three main stages. These stages are:

(i) glycerol purification, (ii) glycerol fermentation, and (iii) propionic acid recovery and

purification. The glycerol purification has been widely discussed throughout this document

(see Chapter 4). On the other hand, for the glycerol fermentation to propionic acid two

strains were identified as the most promissory bacteria available from literature. The first

one is a commercial Propionibacterium acidipropionici which consumes pure glycerol

[125], and the second one is the engineered Propionibacterium acidipropionici ACK-Tet

[126], which is able to consume both pure and crude glycerol as only source of carbon

and energy.

Respect to the recovery and purification methods of organic acids several alternatives

have been evaluated. Some examples are: liquid extraction [127], reverse osmosis [128],

electrodialysis [129], liquid surfactant membrane extraction [130], anion exchange [63],

precipitation and adsorption [131], and reactive liquid-liquid extraction [132]. These

processes were above described for recovering and purifying of carboxylic acids.

Otherwise, Keshav et al [133-144] have studied widely the reactive extraction of propionic

acid from a fermentation broth, and different diluent (e.g., benzene, toluene, hexane, n-

heptane, n-octane, n-dodecane, ethyl acetate, butyl acetate, 1-octanol, 2-octanol, 1-

decanol, 1-dodecanol, petroleum ether, paraffin liquid, MIBK, oleyl alcohol, sunflower oil,

) and extractant (e.g., Tri-n-butylphosphate, tri-n-octylamine, Aliquat 336, and tri-n-

octylphosphine oxide)agents have been analyzed. Besides of the final extraction

performance, kinetic behavior has also been studied [135]. Based on the results reported

by Keshav et al [133-144] for the reactive extraction of propionic acid from the

fermentation broth, it was noticed that the best configuration for this process requires the

use of tri-n-octyl amine (TOA) as extractant agent while the diluent agent must be ethyl


acetate. The concentration that leads to the best performance is 0.686 kmol/m3

and the

extraction temperature is 305 K according to Keshav et al [136].

Here and based on the reviewed literature, a technological scheme for propionic acid

production is proposed, simulated, and economically assessed. First of all, five different

scenarios are considered for the glycerol fermentation stage. Scenario 1 uses a

commercial Propionibacterium acidipropionici strain and glycerol diluted at 20 g/L, while in

the Scenario 2 the fermentation media contains the same strain and glycerol diluted at 50

g/L. For these two scenarios the fermentation process is carry out at 30 °C using pure

glycerol [125]. For Scenario 3, Scenario 4, and Scenario 5 the engineered strain

Propionibacterium acidipropionici ACK-Tet is considered and pure glycerol at 46 g/L is

required in Scenario 3. While crude glycerol at 17 g/L is used in Scenario 4, Scenario 5

considers a completely different configuration for the fermentation process. It is a fibrous-

bed bioreactor packed with immobilized cells and fed with pure glycerol at 41 g/L. For

these three scenarios the fermentation process is carry out at 32 °C [126]. In all cases,

100% of glycerol conversion was reached and the fermentation times reported in each

case were: 120 h (Scenario 1), 150 h (Scenario 2), 280 h (Scenario 3), 160 h (Scenario

4), and 104 h (Scenario 5).

Since the analyzed scenarios differ mainly on the fermentation stage according to Table

8.25, which shows the stoichiometry of the fermentation process for each scenario,

different operational conditions, material and energy requirements, and equipment sizes

are required for each scenario. The principles of the reactive extraction process and the

corresponding back extraction process for recovering carboxylic acids from an aqueous

media were above described.

Table 8.25. Stoichiometry of the fermentation process for each scenario.

Scenario Glycerol Propionic Ac. Acetic Ac. Succinic Ac. Biomass MW 92.09 74.08 60.05 118.09 24.73

1 1 0.9821 0.0925 0.0905 0.3481 2 1 0.7086 0.0785 0.0660 0.2820 3 1 0.6713 0.0368 0.0507 0.3442 4 1 0.8826 0.0537 0.0546 0.1903 5 1 0.7334 0.0414 0.0569 -


Table 8.26. Summary of the main simulation results for prpionic acid production process

Scenario 1

Dilutgly Propaci1 Complex2 Aqueous2 Extract2 Organic5 Propacp

Temperature (K) 303.1 303.1 293.1 283.1 283.4 293.1 379.8 Mass Flow (kg/hr)

Water 27995.5 27995.5 27992.7 27763.1 229.6 8.79 0.018 Glycerol 578.9 0 0 0 0 0 0 Cell mass 0 54.05 0 0 0 0 0 Prop. acid 0 457.3 96.00 88.89 7.11 368.4 367.4 Suc. acid 0 67.18 67.19 62.23 4.96 4.96 4.96 Ac. acid 0 457.3 34.89 32.31 2.58 2.58 0.54 Tri-N 0 0 357.9 26.52 331.4 2056.3 0 Complex 0 0 2075.7 0 2075.7 0 0 Ethyl 0 0 11117.4 2754.03 8363.3 0 0

Scenario 2


Temperature (K) 303.1 303.1 293.1 283.1 283.3 293.1 380 Mass Flow (kg/hr)

Water 11007.5 11007.5 11006.4 10724.5 281.9 6.34 0 Glycerol 578.9 0 0 0 0 0 0 Cell mass 0 43.78 0 0 0 0 0 Prop. acid 0 329.96 69.26 55.48 13.78 274.5 274.0 Suc. acid 0 48.99 49.01 39.20 9.80 9.80 9.80 Ac. acid 0 29.63 29.61 23.72 5.88 5.88 1.20 Tri-N 0 0 838.21 53.76 784.4 2029.0 0 Complex 0 0 1497.6 0 1497.6 0 0 Ethyl 0 0 11117.4 1114.0 10003.4 0 0

Scenario 3


Temperature (K) 305.1 305.1 293.1 283.1 283.3 293.1 380 Mass Flow (kg/hr)

Water 11987.5 11987.5 11986.3 11708.6 277.8 6.00 0 Glycerol 578.9 0 0 0 0 0 0 Cell mass 0 53.44 0 0

0 0

Prop. acid 0 312.6 65.63 53.56 12.07 259.1 258.3 Suc. acid 0 37.63 37.67 30.70 6.97 6.97 6.97 Ac. acid 0 13.89 13.87 11.35 2.58 2.58 0.540 Tri-N 0 0 903.6 59.06 844.6 2023.7 0 Complex 0 0 1418.9 0 1418.9 0 0 Ethyl 0 0 11117.4 1203.8 9913.6 0 0


Scenario 4



Water 12036.6 12036.6 12035.4 11757.2 278.2 7.89 0.018 Glycerol 578.9 0 0 0 0 0 0 Cell mass 0 29.55 0 0 0 0 0 Prop. acid 0 411.0 86.30 70.45 15.85 340.47 339.5 Suc. acid 0 40.5 40.50 33.06 7.44 7.44 7.44 Ac. acid 0 20.27 20.30 16.57 3.72 3.72 0.781 Tri-N 0 0 532.6 34.66 498.0 2047.8 0 Complex 0 0 1865.4 0 1865.4 0 0 Ethyl 0 0 11117.4 1213.7 9903.7 0 0

Scenario 5



Water 13492.5 13492.5 13491.2 13218.4 272.8 6.56 0 Glycerol 578.9 0 0 0 0 0 0 Cell mass 0 53.44 0 0 0 0 0 Prop. acid 0 341.5 71.71 59.93 11.78 281.6 280.8 Suc. acid 0 42.24 42.27 35.31 6.97 6.97 6.97 Ac. acid 0 15.63 15.61 13.03 2.58 2.58 0.54 Tri-N 0 0 794.7 53.05 741.66 2029.4 0 Complex 0 0 1550.4 0 1550.4 0 0 Ethyl 0 0 11117.4 1351.4 9766.0 0 0

Finally, the behavior of the reactive extraction process was analyzed based on its

extraction efficiency, the obtained distribution coefficient and the used loading, but no high

differences were found for the extraction efficiency while the distribution coefficient varied

from 3.8 to 15. Also, it was observed that the global recovery of propionic acid from the

fermentation broth is directly related to the extraction efficiency in the reactive extractive

process as shown in Table 8.27.

Besides of the high recovery levels for propionic acid, the extractan and diluent agents

were almost completely recovered during the downstream process. Thus, low quantities

of both agents must be fed as a fresh stream. As it could be expected, the global yield for

the downstream process agreed with the order found for the fermentation yield and the


propionic acid production (i.e, Scenario 1 > Scenario 4 > Scenario 5 > Scenario 2 >

Scenario 3) with a relative difference between Scenario 1 to Scenario 3 of 29.69%.

Table 8.27. Data representing the behavior of the downstream process for propionic acid

production

Scenario 1

Scenario 2

Scenario 3

Scenario 4

Scenario 5

Reaction-Extraction Process Extraction efficiency (%) 80.56 83.19 82.87 82.85 82.46

Distribution coefficient 8.761 3.790 14.156 12.965 15.032

Loading (Z) 0.844 0.629 0.594 0.781 0.645

Downstream Process Global P.A. recovery (%) 80.34 83.05 82.64 82.61 82.24


EtAc recovery (%) 99.53 99.45 98.51 98.49 98.34

Global Process Global process yield from

Glycerol to PA (%)

0.7605

0.5672

0.5347

0.7027

0.5814

PA: Propionic acid. EtAc: Ethyl acetate.

8.7 Economic assessment

8.7.1 1,3-propanediol production For most industrial processes the cost of raw material represents near to 50 % of the total

production costs, while if raw glycerol is used for the production of 1,3-propanediol by

mean of engineered K. pneumoniae strains this value is lower between 8.2 to 9.2 % of the

total production cost as shown in Table 8.28. The values here obtained do not consider

the transportation costs, since the 1,3-propanediol production process was assumed to be

a biorefinery adjacent the biodiesel production process. Both utilities and capital costs

represent the highest production cost which combined vary between 64.46 % for the

Scenario 3 to 66.99 % for the Scenario 1. Also, it was noticed that capital costs increased

when final concentration of 1,3-propanediol increased.


Table 8.28. Economic results for raw glycerol conversion to 1,3-propanediol: Cost

(USD$/kg) and Share (%)

Scenario 1 Scenario 2 Scenario 3

Raw materials Cost 0.0889 0.0892 0.0887

Share 8.61 8.19 9.20

Utilities Cost 0.2151 0.2245 0.1458

Share 20.83 20.61 15.12

Operating labor Cost 0.0698 0.0746 0.0714

Share 6.76 6.85 7.40

Maintenance Cost 0.0479 0.0503 0.0486

Share 4.64 4.62 5.04

Operating charges Cost 0.0172 0.0187 0.0178

Share 1.66 1.71 1.85

Plant Overhead Cost 0.0577 0.0625 0.0600

Share 5.59 5.73 6.22

G and A cost Cost 0.0650 0.0712 0.0615

Share 6.29 6.54 6.38

Depreciation of capital Cost 0.4767 0.5042 0.4757

Share 46.16 46.28 49.34

Co-products sales Cost -0.0055 -0.0057 -0.0054

Share -0.53 -0.52 -0.56

Gly. Purify. + ferm. Cost 0.304 0.331 0.313

Share 29.40 30.38 32.42

1,3-PD purification Cost 0.729 0.758 0.652

Share 70.60 69.62 67.58

Total Cost 1.033 1.089 0.964

Sale price/production cost 1.710 1.622 1.832

In order to obtain more detailed information about the 1,3-propanediol production from

raw glycerol, the total production cost was divided in two processing sections: (i) Glycerol

purification plus glycerol fermentation and (ii) 1,3-propanediol recovery and purification,

as shown in Table 8.28. The first section (i.e., glycerol purification and fermentation)

represents in all cases between 29.4 to 32.4 %, while the second section (i.e., 1,3-


propanediol purification) is between 67.6 to 70.6 %. The 1,3-propanediol purification cost

was estimated to be between 0.652 to 0.758 USD$/kg of 1,3-propanediol. The increasing

order for the total production is: Scenario 3 < Scenario 1 < Scenario 2, which is not

related to the order found to the 1,3-propanediol production, neither its yield.

Also, the commercial sale price for 1,3-propanediol was compared to the production cost

and it was found that this ratio (i.e., sale price/production cost) is higher than the unity for

all scenarios, and for the Scenario 3 this ratio was 1.832. These results indicate that 1,3-

propanediol production from raw glycerol by mean of engineered K. pneumoniae could be

a profitable alternative for glycerol usage.

8.7.2 Ethanol production The economic assessment for the purification process considers two scenarios. In the first

one, the retired methanol from raw glycerin stream is considered as a waste, and in the

second scenario the methanol is recycled to the transesterification process, which

contains 99 wt % of methanol and 1 wt % of glycerol. Therefore, in the second case

methanol is obtained as a co-product as it was shown in Chapter 4.

Raw glycerol in bioethanol production represented only 30% of the total costs as shown in

Table 8.29. Transportation costs were not considered in the economic assessment since

the purification step was assumed to be adjacent to the biodiesel production process. On

the other hand, utilities and capital costs represent the highest cost on the purification

process (i.e., between 20% and 30%). Also, the final quality of glycerol increases mainly

the utility costs. Purification costs of raw glycerol to pure glycerol at 88 wt % are 0.1574

US$/L (scenario I) and 0.0984 US$/L (scenario II) when the methanol price is considered.

Moreover, when glycerol at 98 wt % is used the Purification costs are 0.1782 US$/L

(scenario I) and 0.1124 US$/L (scenario II). Approximate costs for refining raw glycerol

have been reported to about 0.15 US$/lb or 0.26 US$/L [37], which are higher than the

purification costs obtained here, but near to the obtained purification costs in the scenario

I. Also, purification costs obtained are lower than the sale price of each product, which are

0.28 US$/L for glycerol at 88 wt %, 1.39 US$/L for glycerol at 98 wt % from vegetable oil

and 1.11 US$/L for glycerol at 98 wt % from tallow. Then a decrease in the whole fuel

ethanol production costs from glycerol can be expected due to the purification process.


The economic assessment carried out for the glycerol bioconversion process to fuel

ethanol does not consider the raw material cost because it is involved in the purification

costs. Table 8.29 shows the bioconversion costs (BCCs) obtained using Aspen Icarus.

The lowest BCC was obtained for crude glycerol (88 wt %) when it was diluted at 20 g/L,

since it uses a lower quantity of water than the other two processes. In this way,

equipment size and utilities are modified in each case. On the other hand, when pure

glycerol (98 wt %) is used a higher water quantity is necessary then increasing sizing and

utilities.

Table 8.29. Bioconversion costs (BCCs) for fuel ethanol production form raw glycerol

Item (US$/L)

Crude glycerol (10 g/l)

Share (%)


Share (%)

Pure glycerol (10 g/l)

Share (%)

Utilities 0.0599 31.82 0.0503 29.41 0.0975 41.28

Operating labor 0.0188 9.97 0.0188 10.99 0.0188 7.96

Maintenance and operating charges 0.0205 10.89 0.0154 9.02 0.0193 8.17

Plant overhead and general and administrative costs 0.0266 14.14 0.0309 18.05 0.0410 17.37

Capital depreciation 0.0625 33.18 0.0556 32.54 0.0596 25.23

Product production cost (US$/L) 0.1883 100.00 0.1710 100.00 0.2361 100.00

Finally, global production costs (GPCs) for raw glycerol bioconversion to fuel ethanol are

obtained by adding the purification costs and the BBCs, like shown in Table 8.30. In all

cases the purification costs are near 35 % and the BCCs are near 65 %. Furthermore,

these obtained GPCs from crude glycerol are lower than those reported by Quintero et al

[17] for fuel ethanol production from corn (0.3381 US$/L), where using crude glycerol at

10 g/L and 20 g/L could represent a saving of 15 % and 20 %, respectively. The obtained

GPCs are higher than those reported by Quintero et al [17] for fuel ethanol production

from sugar cane (0.2153 US$/L). Nevertheless, these obtained GPCs are lower than the

international prices for fuel ethanol ranging from (0.4552 US$/L [145] to 0.6057 US$/L

[146]). The


Although a rigorous analysis of ethanol fuel market and its prices should be carried out,

the obtained results indicate that the production process of ethanol fuel from raw glycerol

using E. coli can be as profitable as those using sugar cane or corn as feedstocks.

Table 8.30. Global production costs (GPCs) for fuel ethanol production from raw glycerol.

Costs


Share (%)


Share (%)

Pure glycerol (10 g/l)

Share (%)

Purification Costs 0.0984 34.32 0.0984 36.53 0.1124 32.26

Bioconversion Costs 0.1883 65.68 0.1710 63.47 0.2361 67.74

Global Costs 0.2867 100.00 0.2694 100.00 0.3485 100.00

On the other hand, for the sustainable production of biodiesel from oil palm, total

production cost of ethanol from crude glycerol was estimated in 0.2694 US$/L, but this

result considers the buying of crude glycerol. This cost is lower than these reported for

fuel ethanol production from corn (0.3381 US$/L) and higher that these reported from

sugar cane (0.2153 US$/L) as showed Quintero et al [17].

Total production cost of ethanol from lignocellulosic biomass has been calculated by

McAloon et al. [147] was 0,396 US$/L, which is lower than the current international price

of fuel ethanol, reported by ICISpricing [145] as 0.4552 US$/L. On the other hand, total

production cost of biodiesel from oil palm under above described operation conditions

was estimated in 0,7971 US$/L (results are not shown), where its sale price ranged

between 0,76 – 0,86 US$/L (ICISpricing, 2010) [145].

Total production cost of biodiesel from oil palm in this integrated production process was

estimated in 0,8803 US$/L, which is higher than biodiesel production cost from palm oil.

This cost has been discriminated by raw material, services, operative and administrative

costs, depreciation and co-products credit, as shows the Table 8.31. Also, column 2

shows the share for each item.


Table 8.31. Discriminated costs for integrated biodiesel and raw-ethanol production from

oil palm.

Item Cost (US$/L) Share (%)

Raw materials 0,3376 35,69

Service fluids 0,2105 22,25

Labour 0,0083 0,87

Maintenance 0,0621 6,57

Operating charges 0,0301 3,18

General and administrative costs

0,0750 7,92

Capital depreciation 0,2225 23,52

Co-products credit -0,0657 -6,95

Total production cost 0,8803 100,00

This technology is promising from an environmental viewpoint because of petrochemical

methanol is not required and also in the case of oil palm industry an independence from

feedstock can be achieved. On the other hand, regarding crude glycerin, traditional

biodiesel industries have had only two options, the first one is its sale as low-quality

glycerin with a rather low price; the second is its refining until either, industrial or USP

grades, where the involved purification cost have been previously estimated as 0.20 and

0.27 US$/L, respectively Posada and Cardona [37]. These values agree with the

purification cost reported by Johnson and Taconi [146], which is 0.26 US$/L. Although,

biodiesel production cost by this technology indicates that the process is not economically

viable, a further analysis of process design which considers a higher detail level as:

energy cogeneration, heat integration, and the use of new technologies, among others,

could help to improve the process efficiency which would be reflected directly in the

production cost of biodiesel.

Described technology is a perfect representation of the biorefinery concept applied to the

biofuels production, in which oil palm is used as a sole feedstock for producing mainly

biodiesel, and ethanol as an important co-product which is widely used by the transport

sector. Therefore, it is expected that this integrated technology for biodiesel production

which uses efficiently the lignocellulosic residues and crude glycerol to produce the


required ethanol as feedstock, can be regarded as a valuable opportunity for biofuels

production. Moreover, the two proposed ways for ethanol production should eliminate the

need of using other feedstock which is important as foods, such as: sugar cane, sugar

beet, corn, wheat, sorghum, etc.

8.7.3 PHB production The total production costs of PHB at 99.9 wt % from raw glycerol are shown in Table 8.32.

This table was obtained using glycerol at 88 or 98 wt % in the fermentation stage. Also,

the three downstream processes were compared. In all cases the cost for raw material

represents the glycerol purification cost. In this way, the glycerol purification process

represents only between 4.8 to 5.6 % of the total PHB production cost when glycerol at 88

wt % is used and this values increased between 6.3 to 7.7 % when glycerol at 98 wt % is

used.

Table 8.32. Total PHB production costs from crude glycerol through raw glycerol (88 wt

%) and pure glycerol (98 wt %)

Item

Cost (US$/kg) and Share (%)

*DSP I *DSP II *DSP III 88 wt % 98 wt % 88 wt % 98 wt % 88 wt % 98 wt %

Raw material

Cost 0.118 0.149 0.118 0.149 0.118 0.149

Share 5.6 7.71 4.84 6.27 5.42 7.07

Utilities

Cost 0.7787 0.658 0.9809 0.953 0.8678 0.841

Share 36.94 33.96 40.23 39.97 39.87 39.77

Operating labor

Cost 0.0752 0.066 0.0752 0.068 0.0752 0.064

Share 3.57 3.41 3.08 2.85 3.45 3.03

Maintenance and operating charges

Cost 0.2296 0.21 0.2402 0.236 0.2451 0.223

Share 10.89 10.84 9.85 9.9 11.26 10.55

Plant overhead, general and administrative costs

Cost 0.2128 0.186 0.2287 0.218 0.2091 0.191

Share 10.09 9.6 9.38 9.14 9.61 9.03

Depreciation of capital

Cost 0.6938 0.668 0.7951 0.76 0.6616 0.646

Share 32.91 34.48 32.61 31.87 30.39 30.55

Total

Cost 2.108 1.937 2.438 2.384 2.1767 2.114 Share 100 100 100 100 100 100

*DSP: Downstream process


In general terms the lower production costs are obtained when glycerol at 98 wt % is used

in the fermentation process, due to the higher PHB yield in the fermentation stage and the

lower energy requirements in the spray drying process. On the other hand, the lower

glycerol purification cost obtained for glycerol at 88 wt % generates a lower share in the

total PHB production costs compared with glycerol at 98 wt %. In both cases (88 and 98

wt %) the higher value in the total production cost was obtained for the Downstream

Process II, which uses a solvent extraction stage. This extraction requires heating the

expensive solvent DES up to 110 °C which increase the utility costs.

The total PHB production costs obtained are between 2.11 and 2.44 US$/Kg when

glycerol at 88 wt % is used in the fermentation process, and between 1.94 and 2.38

US$/Kg when glycerol at 98 wt % is used. These production costs are near to the lower

sale prices reported in the literature (involving profits) and were obtained from other

substrates (see Table 8.33). In economical terms the best technological scheme for PHB

production from crude glycerol includes three important features as follows: i) purification

of crude glycerol up to 98 wt %, ii) a two continuous fermentation stages and finally iii) the

PHB recovering performed with the Downstream Process I, which is similar to the

BIOPOL flow sheet [45-46].

Table 8.33. Main producers of PHA in the world

PHB: Product name

Company Substrate Price (US$/kg) Production (t/y)

Biomer: P(3HB) Biotechnoly Co. Germany

Small scale production [148]

25 (2003) 3.75-6.25 (2010)

50 (2003)

Biocycle: P(3HB) PHB industrial S/A company, Brazil

Sugar cane[149]

12.5-15 (2003) 3.12-3.75 (2010)

1400(2003) 30-60000(2010)

Biogreen: P(3HB)

Mitsubishi GAS Chemical, Japan

Methanol [148]

2.75 (2010)

-

P(3HB) Metabolix, USA (BASF, ADM)

Corn, sugar [149]

- -

Pure glycerol has a higher cost than raw glycerol, but the yield of PHB is higher when

pure glycerol is used as feedstock. Moreover, each downstream process should be

adjusted to the conditions of the fermentation broth. Thus, using the blocks process


analysis (glycerol purification, fermentation, and PHB isolation and purification) the best

technological scheme for the production of PHB was found.

In all the cases assessed in this study the total production costs (see Table 8.32) were

lower than those reported in the literature (see Table 8.33) using feedstocks different to

glycerol. Generally, it has been suggested that the higher share to the total PHB

production cost is the substrate cost which is up to 45%. Meanwhile, if glycerol is used as

feedstock this share is below 8 %. These results indicate that using crude glycerol as

feedstock to produce PHB could be a profitable alternative to develop biorefineries in the

biodiesel industry.

8.7.4 D-Lactic acid production The five scenarios were economically compared as shown in Table 8.34. The total

production cost was discriminated, in general terms, by raw materials, services,

operatives, depreciation, and products and co-product sales.

If raw glycerol is used for the D-lactic acid production by mean of engineered E. coli

strains this value is lower than 8 % of the total production cost. The values here obtained

do not consider the transportation costs, since the D-lactic acid production process was

assumed to be a biorefinery adjacent the biodiesel production process. Both utilities and

capital costs represent the highest production cost which combined vary between 69.04

% for the Scenario 3 to 73.12 % for the Scenario 1. Also, it was noticed that utilities costs

increased when glycerol concentration decreased, which makes sense due to the higher

requirement of both water and heat for the fermentation stage. The opposite behavior

occurred for the capital costs.

In order to obtain more detailed information about the D-lactic acid production from raw

glycerol, the total production cost was divided in two processing sections: (i) Glycerol

purification plus glycerol fermentation and (ii) D-lactic acid recovery and purification, as

shown in Table 8.32. The first section (i.e., glycerol purification and fermentation)

represents in all cases between 21 to 26 %, while the second section (i.e., D-lactic acid

purification) is between 74 to 79 %. The D-lactic acid purification cost was estimated to be

between 0.789 to 0.915 USD$/kg of L.A. Although these values are lower than that ones


reported by Joglekar et al. [55] for a similar downstream process (see: Table 8.15, Route

1), the esterification and hydrolysis by reactive distillation processes were not here

considered. Also, a rigorous economic analysis was carried out meanwhile the purification

costs obtained by Joglekar et al. [55] were based only on reported literature.

Table 8.34. Economic results for raw glycerol conversion to D-lactic acid: Cost (USD$/kg)

and Share (%)


Raw materials Cost 0.081 0.0812 0.0789 0.0791 0.0252

Share 6.93 6.78 7.16 7.79 2.43

Utilities Cost 0.4556 0.3932 0.2956 0.2844 0.2736

Share 38.95 32.85 26.83 28.02 26.35

Operating labor Cost 0.0542 0.0583 0.0565 0.0565 0.0581

Share 4.63 4.87 5.13 5.57 5.60

Maintenance Cost 0.0488 0.0628 0.0659 0.0579 0.069

Share 4.17 5.25 5.98 5.70 6.65

Operating charges Cost 0.0136 0.0146 0.0141 0.0141 0.0146

Share 1.16 1.22 1.28 1.39 1.41

Plant Overhead Cost 0.0515 0.0606 0.0612 0.0572 0.0635

Share 4.40 5.06 5.56 5.64 6.12

G and A cost Cost 0.0707 0.0788 0.0699 0.0681 0.0703

Share 6.04 6.58 6.34 6.71 6.77

Depreciation of capital Cost 0.3997 0.453 0.465 0.4468 0.4696

Share 34.17 37.85 42.21 44.02 45.23

Co-products sales Cost -0.0055 -0.0057 -0.0054 -0.0492 -0.0056

Share -0.47 -0.48 -0.49 -4.85 -0.54

Gly. Purify. + ferm. Cost 0.304 0.282 0.274 0.226 0.219

Share 25.98 23.56 24.86 22.27 21.10

L.A. purification Cost 0.866 0.915 0.828 0.789 0.819

Share 74.02 76.44 75.14 77.73 78.90

Total Cost 1.1696 1.1968 1.1017 1.0149 1.0383

Sale price/production cost 1.327 1.297 1.409 1.530 1.495


High differences can be observed for the total production cost of D-lactic acid from raw

glycerol, which range from 1.0149 to 1.1968 USD$/Kg of L.A. The increasing order for the

total production cost is: Scenario 4 < Scenario 5 < Scenario 3 < Scenario 1 < Scenario 2,

which is completely related to the order found to the D-Lactic acid production. Also, the

commercial sale price for D-lactic acid was compared to the production cost and it was

found that this ratio (i.e., sale price/production cost) is higher than the unity for all

scenarios, and for the Scenario 4 this ratio was 1.53. These results indicate that D-lactic

acid production from raw glycerol by mean of engineered E. coli could be a profitable

alternative for glycerol usage.

The lowest total production cost of D-lactic acid from raw glycerol was obtained for the

Scenario 4 in which three fermentative advantages are joined simultaneously: (i) use of

crude glycerol (85 wt %), (ii) high glycerol concentration in the fermentation media (40 g/l),

and (iii) total consumption of glycerol. Thus, the economical successful of a

biotechnological scheme depends highly on the fermentation stage and the above

described three characteristics could lead to high process performance. In this order of

ideas, one of the main purposes of metabolic engineering should be to develop specific

strains able to consume completely raw substrates at high concentrations.

8.7.5 Succinic acid production The five scenarios were economically assessed and the results are shown in Table 8.34.

These production costs were discriminated by raw materials, services, operatives,

depreciation, and products and co-product sales. The values here obtained do not

consider the transportation costs, since the succinic acid production process was

assumed to be a biorefinery adjacent the biodiesel production process.

In all cases the raw material costs represent only between 2.8 to 3.91 %, while the utilities

are the highest share of the total production costs with a weight between 65.9 to 70.12 %

followed by the capital costs with a share among 11.23 to 14.55 %.


Table 8.35. Economic results for raw glycerol conversion to succinic acid: Cost (USD$/kg)

and Share (%)



Share 3.91 3.00 3.41 3.33 2.79

Utilities Cost 1.5673 2.0850 1.7877 1.8644 2.1486

Share 67.60 68.95 68.79 70.12 65.93


Share 2.51 1.94 2.14 1.89 1.79


Share 1.58 1.82 1.72 1.72 2.05


Share 0.63 0.49 0.53 0.47 0.45


Share 2.05 1.88 1.93 1.81 1.92

G and A cost Cost 0.1720 0.2882 0.2547 0.2561 0.4072

Share 7.42 9.53 9.80 9.63 12.50


Share 14.55 12.58 11.92 11.23 12.74


Share -0.24 -0.18 -0.23 -0.21 -0.17


Share 13.39 10.60 10.92 9.46 9.52

S.A. purification Cost 2.008 2.703 2.315 2.407 2.949

Share 86.61 89.40 89.08 90.54 90.48

Total Cost 2.319 3.024 2.599 2.659 3.259


Also, it was noticed that utilities costs increased with both the global fermentation yield

(see Table 8.20) and the low concentration of glycerol in the fermentation media. This

behavior is due to both the inefficient usage of glycerol during the fermentation process

and the higher requirement of serves.


The total production costs were divided in two processing sections: (i) Glycerol purification

plus glycerol fermentation and (ii) succinic acid recovery and purification, as shown in

Table 8.32. Thus, more detailed information about the succincic acid production from raw

glycerol can be obtained. In this way, glycerol purification and fermentation represent, in

all cases, lower than 13.4 % indicating that most of the total production cost of succinic

acid from raw glycerol are in the recovery and purification processes (> 86.5 %).

Regarding to the succinic acid purification costs, high differences were found among the

five scenarios since this value varies from 2.008 to 2.949 USD$/Kg of S.A, being its

relative difference almost 50%. These differences found in the purification costs are

mainly related to the complex behavior of the fermentation process, due to variables such

as glycerol concentration, glycerol purity, glycerol consumption, and yield to succinic acid

affects the performance of the global process. Thus, the increasing order for succinic acid

recovery and purification costs and for total production costs of succinic acid from raw

glycerol are the same, it is Scenario 1 < Scenario 3 < Scenario 4 < Scenario 2 < Scenario

5.

On the other hand, the commercial sale price for succinc acid was compared to its

production cost and it was found that this ratio (i.e., sale price/production cost) was only

higher than the unity for the Scenario 1, but this value is still too close to the unity. Thus, it

can be stated that the succinic acid production from glycerol still requires performing high

effort in order to make this process profitable since the its production costs is high yet.

8.7.6 Propionic acid production The five analyzed scenarios were economically assessed as shown in Table 8.36 and the

production costs results were discriminated by raw materials, services, operatives,

depreciation, and products and co-product sales. The values here obtained do not

consider the transportation costs, since the propionic acid production process was

assumed to be a biorefinery adjacent the biodiesel production process.

Low raw material costs were obtained in all cases when raw glycerol is used, representing

only around 5 % of the total production costs. But both utilities (between 45.8 to 58.9 %)

and capital (between 20.8 to 29.9 %) costs represent the highest production cost which


combined vary between 75.76 % for the Scenario 3 to 79.69 % for the Scenario 1 and the

Scenario 5. Also, it was found that the utilities costs depend mainly on the glycerol

concentration in the fermentation media.

Table 8.36. Economic results for raw glycerol conversion to propionic acid: Cost

(USD$/kg) and Share (%).



Share 4.41 5.05 4.71 4.38 4.59

Utilities Cost 1.1355 0.8192 0.8275 0.9709 0.9497

Share 58.89 48.67 45.85 49.99 51.22


Share 2.91 3.71 3.91 3.83 3.73


Share 1.78 3.87 4.34 4.33 4.11


Share 0.73 0.93 0.98 0.96 0.93


Share 2.34 3.79 4.12 4.08 3.92

G and A cost Cost 0.1625 0.1161 0.1171 0.0636 0.0618

Share 8.43 6.90 6.49 3.27 3.33


Share 20.80 27.41 29.91 29.43 28.47


Share -0.28 -0.33 -0.30 -0.28 -0.30


Share 16.07 18.46 17.20 15.23 16.74

L.A. purification Cost 1.618 1.372 1.495 1.647 1.544

Share 83.93 81.54 82.80 84.77 83.26

Total Cost 1.928 1.683 1.805 1.942 1.854



In addition, the total production cost was divided in two processing sections: (i) glycerol

purification plus glycerol fermentation and (ii) propionic acid recovery and purification, as

shown in Table 8.36. The glycerol purification and glycerol fermentation stages only

represented between 15.23 to 18.46 %, which indicates that most efforts are required for

the downstream process since a high share of the total production costs (between 81.54

to 84.77 %) is consumed for the propionic acid purification.

Total production costs of propionic acid from raw glycerol range from 1.683 to 1.942

USD$/Kg of P.A., which represents a relative difference of 15.4 %. Being the increasing

order for the total production costs: Scenario 2 < Scenario 3 < Scenario 5 < Scenario 1 <

Scenario 4.

The ratio between the sale price to the production cost was calculated for the five

analyzed scenarios, and this ratio was higher than the unity only for the Scenario 2,

Scenario 3, and Scenario 5. Thus, it is stated that this process could be profitable only

when high concentration of glycerol is used in the fermentation media and high yield to

propionic acid is obtained.

8.8 Conclusions In the past, important efforts have been made in order to introduce the biotechnological

production of 1,3-propanediol from glycerol to the industry. However, research tendencies

were focused on microorganism development and on the analysis of some specific

process conditions. On the other hand, the drastic increment in the use of biodiesel has

caused an oversupply of glycerol in the market. Thus, in order to suitably use the glycerol

obtained from the biodiesel industry, the mass production of 1,3-propanediol from glycerol

needs additional analysis. Here, a complete technological scheme for its production has

been not only proposed by also assessed. Even more, the fermentation stage was

optimized by three different ways and as result it was possible to assess economically

three scenarios where the third one had the lowest production cost. The relative

differences respect to the first and the second scenarios were 7.2% and 13.0 %. The

obtained results provide enough information to understand the different possibilities for

process intensification using this technology and also to compare it with other new

industrial alternatives for the utilization of glycerol as a raw material.


Due to the low cost of raw glycerol, methanol recovery from glycerol implies low

purification costs. Meanwhile, the three possibilities assessed for glycerol bioconversion

to ethanol showed that the global production costs of fuel ethanol from raw glycerol are

lower than the commercial price of fuel ethanol. These facts show the potential for raw

glycerol bioconversion to fuel ethanol using E. coli. Also, the comparison carried out with

a previous paper (which considers the fuel ethanol production from sugarcane and corn in

the Colombian case) shows that the global production costs of fuel ethanol from raw

glycerol can be as profitable as the production of fuel ethanol from conventional raw

materials as sugarcane. The latter is a completely developed industry in Colombia.

Production of ethanol from glycerol is presented as an alternative which uses a renewable

resource that does not generate direct or indirect competition with the food industry,

property that do have the raw materials used in commercial (maize, cassava Beets, sugar

cane, molasses, etc.). The techno-economic feasibility of this process was demonstrated

through simulation and evaluation of processes and found that the cost of production is

similar to the process established commercially in Colombia from sugar cane. The

environmental performance of this process was not the best of those obtained for the

different raw materials, but significant higher compared with the process benefits from

corn.

Biodiesel and ethanol can be jointly produced using oil palm as sole source by mean of

processes integration, such as the biodiesel production with the ethanol production from

two feedstocks: lignocellulosic residues (empty fruit bunches and palm press fiber

produced during) and crude glycerol. Thus, alcohol is completely self-supplied by the

integrated process and low quantities of wastes are produced without any global

production of glycerol. Economical evaluation showed a higher biodiesel production cost

for the integrated process than the traditional biodiesel production process which uses

ethanol and palm oil as feedstocks. But the first one is a promising technology available to

build an autonomous biodiesel production plant with low waste levels. This process must

be economically improved by further analysis from a process design view point, based on

process simulation which showed be a powerful tool for performing processes integration.

Three technological schemes to produce PHB from crude glycerol were analyzed under

two fermentation conditions (i.e., using glycerol at 88 wt % and 98 wt %). In this work it

was found that it is better to use pure glycerol as feedstock for the production of PHB than


raw glycerol. This phenomenon is explained by the fact that the higher PHB yield reduces

the utility costs in the downstream process. The results shown here are important for the

industrial production of PHB using glycerol as a raw material. Currently, in the biodiesel

production the total profitability of any new project could be determined by the right use of

glycerol as a massive by-product. The proposed strategy to use pure glycerol as

substrate can be understood as a very interesting alternative since the final composition

in the glycerol streams depends on the source of the feedstock used for biodiesel

production. Also, most of the biotechnological alternatives to produce added value

compounds from glycerol are sensitive to contaminants in the raw material. Thus, several

technical and economical advantages as well as a more stable production of PHB are

obtained when a standardized raw material as pure glycerol is used.

Usage of raw glycerol, engineered Escherichia coli strains, and processes integration for

the production of optically pure D-lactic acid is an important alternative to transform the

by-produced glycerol during the biodiesel synthesis. Although five different configurations

for the fermentation stage were considered, in all cases the total production costs were

lower than its sale price. Thus, the whole process scheme for D-lactic acid production

could be considered as potentially profitable design. Also, it was found that the combined

effect of both high glycerol concentration and use of low quality glycerol in the

fermentation media, lead to the best economic performance. The results shown here are

important for the industrial production of D-lactic acid using glycerol as a raw material. On

the other hand, in the biodiesel production the total profitability of any new project could

be determined by the right use of glycerol as a massive raw material.

Otherwise, for succinic acid and propionic acid production results showed that most

efforts are required in order to reduce the downstream process since the total production

costs were too close to the sale price of these carboxylic acids.

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9. Experimental Setup for Glycerol Fermentation to PHB

Polyhydroxyalkanoates (PHAs) have been recognized as good substitutes for the non-

biodegradable petrochemically produced polymers. However, their high current

production cost limits their industrial applications. Carbon source can be represent

between 25% to 45% of the total PHB production costs. Glycerol, a by-product from the

biodiesel industry, can be used as primary carbon source for cell growth and PHB

synthesis and it is an interesting alternative for to increase the PHB feasibility economic

process. Currently, PHB is produced in an industrial scale using Gram negative bacteria.

Nevertheless, Gram-negative organisms contain lipopolysaccharides (LPS) which

copurify with the PHAs and induces a strong immunogenic reaction. Gram-positive

bacteria lack LPS and are hence potentially better sources of PHAs when used for

biomédical purposes. In this work, the conditions and capability of poly (β-

hydroxybutyrate) (PHB) production by a Bacillus megaterium (Gram Positive bacteria

isolated from superficial sediments of Bahía Blanca Estuary (Buenos Aires, Argentina))

using glycerol as only carbon source were studied. This microorganism was adapted

and tested at different initial glycerol concentrations and compared with other substrates

as glucose. Aspen Plus and Aspen Icarus were used for the processes simulation and for

the economic assessment, respectively.

9.1 Generalities Polyhydroxyalcanoates are attractive substitute biopolymers for conventional

petrochemical plastics which have similar physical properties to thermoplastics and

elastomers. PHAs are homo or heteropolyesters and can be synthesized and stored

intracellularly by many bacterias in the form of granules and can account for up to 80% of

the total bacterial dry weight [1] [2]. They can be produced from renewable resources

through a fermentation process under restricted growth conditions for nitrogen,


phosphorus, sulfurs and/or oxygen in the presence of an excess carbon source [3]; PHAs

are also completely biodegraded and biocompatible [1], [2], [3], [4].

Polyhydorxybutyrate (PHB) was the first type of PHAs discovered and the most widely

studied. PHB has similar mechanical properties to conventional plastics like

polypropylene or polyethylene, but its production cost are higher than the petrochemical

plastic. The PHB production cost depends on the microorganism (yield and productivity),

carbon and nitrogen sources (substrates), fermentation conditions (temperature,

aeration), recovery and purification of the PHB. Carbon source could represent between

25 to 45 % of the total production costs [4][5].

Many researches have been developed in order to find cheaper carbon sources. Agro-

industrial wastes are attractive candidates because they have desired characteristics:

such as: low prices and high availability. Moreover, when these wastes are used an

environmental problem is avoided.

There are different kinds of microorganisms able to produce PHB from diverse

agroindustrial wastes. Substrates such as whey, lignocellulosic materials and glycerol

from biodiesel have been studied (Table 9.1).

Glycerol is by-generated during the biodiesel production. With every 100 lbs of biodiesel

produced by the transesterification of vegetable oils or animal fats, 10 lbs of crude

glycerol are generated (10 wt %) [16]. Although pure glycerol is an important industrial

feedstock used in foods, drugs, cosmetics, pharmaceuticals, pulp and paper, leather,

textile and tobacco industries, the growth on biodiesel industry has carried out a glycerol

surplus and its consequent price has decreased. Thus, the economy of biodiesel industry

has been directly affected.

Chemical and biological transformations have been analyzed in order to convert glycerol

to added-valuable products. Biological conversion offers the opportunity to synthesize a

large array of products and functionalities. Glycerol can be used such carbon source in

microbiological process substituting sugars owing to the highly reduced nature of carbon

atoms. Different works have shown that some strains of microorganisms can produce

PHB using crude glycerol as carbon and energy source. This microorganisms can be wild

9. Experimental Setup 185

strains such as Cupriavidus necator [12], Methylobacterium rhodesianum [13] or

recombinant microorganism such as E. coli recombinant [15-16]. In the group of PHB

producer bacteria there are both Gram positive and Gram negative strains. Currently,

PHB is produced in an industrial scale using Gram negative bacteria. Nevertheless,

Gram-negative organisms contain lipopolysaccharides (LPS) which copurify with the

PHAs [17] and induces a strong immunogenic reaction. Therefore it is undesirable for

biomedical applications. Those purification processes increase the PHB production costs.

Gram-positive bacteria lack LPS and are hence potentially better sources of PHAs when

used for biomedical purposes [18].

Table 9.1. Some microorganisms PHB producer from different agroindustrial wastes.

Agroindustrial Waste Microorganism Productivity

(gPHB l-1h-1 Reference )

Whey

Methylobacterium sp. ZP24

1.18 [5]

Pseudomonas hydrogenovora

0.18 [6]

Thermus thermophilus HB8

[7]

Recombinant Escherichia coli

0.90 [8]

Lign

ocel

lusi

c

Cane bagasse

Burkholderia sacchari IPT 101

0.11 [9]

Burkholderia cepacia IPT 048

0.09 [9]

Ralstonia eutropha [10] Xilosa + Glucose

B. cepacia ATCC 17759 0.47 [11]

Glycerol

Cupriavidus necator DSM 545

1.51 [12]

Methylobacterium rhodesienum MB 126

0.26 [13]

Osmophilic organism 0.05 [14] Escherichia coli CT1061 0.18 [15]

Bacillus megaterium is Gram positive, strict aerobic, non motile, rod shaped, spore

forming, citrate positive, bacteria hydrolyzing gelatin and casein [19]. We continue the

study of B. megaterium isolated from superficial sediments of Bahia Blanca Estuary


(Buenos Aires, Argentina) after to check in other work [20] that this microorganism is a

PHB producer from glucose. In this paper, we report the total adaptation of B. megaterium

to glycerol as carbon source.

9.2 Materials and Methods

9.2.1 Bacterial strain and its maintenance The Bacillus megaterium was isolated from superficial sediments of Bahia Blanca Estuary

(Buenos Aires, Argentina) and characterized as a PHB producer in the presence of

excess carbon source with the restriction of nitrogen. This strain was used for the current

study. Stock cultures were grown at 33 °C in nutrient broth and maintained at 4°C after

growth on nutrient agar during the activation level. After adaptation to glycerol as sole

carbon source, the stock culture was maintained at 4°C after growth on formulated agar

with glycerol. B. megaterium cells were stored at -80 °C in 2 ml cryovials containing 300

µl of glycerol and 700 µl of a previously prepared growth liquid culture.

9.2.2 Culture medium The seeding medium was prepared with the following concentrations: (NH4)2SO4 , 1g/l;

KH2PO4 , 1.5 g/l; Na2HPO4 , 9 g/l; MgSO4 · 7H2O, 0.2g/l; and 1 ml of trace element

solutions composed by: FeSO4 · 7H2O, 10 g/l; ZnSO4 · 7H2O, 2.25 g/l; CuSO4 · 5 H2O,

1g/l; MnSO4 · 4H2O, 0.5; CaCl2 · 2H2O, 2 g/l; H3BO4 , 0.23; (NH4)2Mo7O24 , 0.2 g/l; and

HCl, 10ml. The carbon sources analyzed are both glycerol and glucose. The carbon

source and MgSO4 .7H2

9.2.3 Microorganism’s adaptation and culture conditions

O were autoclaved separately and added aseptically to the

medium after cooling.

B. megaterium cells were used to inoculate nutrient agar plates supplemented with 20 g/l

of glycerol previously incubated at 30°C until growth and then stored at 4ºC. Single

colonies from the grown nutrient agar plates were inoculated in 100ml erlenmeyer flasks

containing 10 ml of sterile nutrient broth medium supplemented with 20 g/l of glycerol and

incubated at 30°C in an orbital shaker at 200 rpm during a period of 2-3 days. 1 ml of

these cultures was again transferred to 10 ml of seeding medium with glycerol.

Successive subculturing was performed 2-3 times, to assure a good adaptation of cell


growth on glycerol as carbon source. The next step was the seed of the culture in the

formulated medium. 1 ml of culture growth in nutrient broth supplemented with 20 g/l of

glycerol was transferred to 10 ml of formulated medium with 20 g/l of glycerol. Successive

subculturing was performed 2-3 times, to assure a good adaptation of cell growth

formulated medium and glycerol. Those tests were incubated at 30°C in an orbital shaker

at 200 rpm during a period of 2-3 days.

9.2.4 Batch cultivations The PHB fermentation was carried out for 36 h in a 3.7 liters Lab Fermenter

(Bioengineering, Switzerland) by using a grown formulated medium. Two different carbon

sources are studied: glucose and glycerol. For both, the formulated medium and

fermenter are the same. The fermentations with glycerol were carried out at 30°C and

33°C and 200rpm. Two different initial glycerol concentrations are used 20 and 50 g/l. The

fermentation with glucose was carried out at 33°C and 200 rpm. The initial glucose

concentration was of 20g/l. The culture volume was of 1.5 l for all batch fermentations.

9.3 Analytical Methods

9.3.1 Biomass The biomass of the culture was determined using the optical density measurement at

600nm (Spectronicspectrophotometer, Thermo SCIENTIFIC GENESYS 20) and by

gravimetric.

9.3.2 PHB extraction After fermentation, the cells were harvested by centrifugation at 18°C and 6.000 rpm for

20 min and then the intracellular PHB was extracted by using the Chloroform–

hypochlorite dispersion extraction. The dispersion media contains 50ml of chloroform and

50ml of a diluted (30 wt %) sodium hypochlorite solution in water, in an orbital shaker at

100 rpm. The cell powder was treated at 38°C for 1 h. The mixture obtained was then

centrifuged at 4000 rpm for 10 min, which resulted in three separate phases. PHB was

recovered from the bottom phase that contains PHB dissolved in chloroform. PHB is

precipitated using 10 volumes of ice-cold methanol [21].


9.3.3 PHB quantification Dried biomass is used for methanolysis of monomers according to the method described

by Braunegg et al. [22] and modified by Lageveen al. [23]. Approximately 10 mg of cells

mass was reacted in a small screw-cap test tube with a solution containing 1 ml of

chloroform, 0.85 mL of methanol, and 0.15 mL of sulfuric acid for 140 min at 100 °C. After

reaction, 0.5 mL of distilled water was added and the test tube was shaken vigorously for

1 min. After phase separation, the organic phase (bottom layer) was removed and

transferred to a small screw-cap glass vial. 50 μl from this organic phase were taken and

added to a test tube and injected in the GC-MS. And an Agilent Technologies 6850 series

II gas chromatograph was used. The gas chromatograph was equipped with a HP-5MS

capillary column (25 m length, 0.32 mm internal diameter). Helium (5cm/min) was used as

the carrier gas. Injector and detector were operated at 230 °C and 275 °C, respectively. A

temperature program was used for efficient separation of the esters (120 °C for 5min,

temperature ramp of 8 °C per min, 180 °C during 12 min). An Agilent Technologies 5975B

mass spectrometer was used.

9.3.4 Glycerol quantification Glycerol concentration was determined off-line by HPLC (Hitachi LaChrom Elite)

equipped with an auto sampler (Hitachi LaChrom Elite L-2200), a Bio-Rad Aminex

Fermentation Monitoring Column (150 mm x 7.8 mm), a column oven (Hitachi LaChrom

Elite L-2300), a HPLC pump (Hitachi LaChrom Elite L-2130) and a Hitachi LaChrom Elite

L-2490 refraction index detector. Injection volume was 20 ml and elution was achieved

using a 50 mM solution of H2SO4. The column was kept at 65°C and the pump was

operated at a flow rate of 0.8 ml min-1

9.3.5 PHB characterization

.

Fourier transform infrared spectroscopy (FTIR): Freeze-dried, precipitated PHB from B.

cereus SPV was used to prepare KBr discs (sample:KBr, 1:100). An FTIR spectrum

1720X spectrometer (Perkin Elmer, USA) was used under the following conditions:

spectral range, 4000–400 cm−1; window material, CsI; 16 scans; resolution 4 cm−1; the

detector was a temperature-stabilized, coated FRDTGS detector [24].


9.4 Results and discussion Glycerol was used as only carbon source and we studied the productivity and yield of the

microorganism. The fermentation profile of glycerol to PHB and biomass, in formulated

medium by B. megaterium are shown in Figures 9.1 and Figure 9.2. Each one of

fermentation was carried out with different initial glycerol concentrations (20g/l and 50 g/l,

respectively). Both fermentations were carried out at 30°C and 200 rpm, in the 3.7 liters

Lab Fermenter (Bioengineering, Switzerland). The culture volume was of 1.5 l. Final

concentration of PHB was 1.116 g/l with a productivity of 0.0248 g/l*h when glycerol at 20

g/l was used, while the final concentration of PHB was 2.356 g/l with a productivity of

0.7365 g/l*h when glycerol at 50 g/l was used. Therefore, the best initial glycerol

concentration is the 50 g/l.

Figure 9.1. Accumulation profile I of Bacillus megaterium cultivated in the 3.7 liters Lab

Fermenter (Bioengineering, Switzerland) with a culture volume of 1.5 l, initial glycerol

concentration of 20 g/l at 30ºC and 200 rpm. Biomass, PHB.

Then, the influence of temperature on the fermentation process was analyzed. The

temperature range for the growth of Bacillus was found to be from 25 to 45°C [17]. The

temperatures analyzed in this work were 30°C and 33°C. Both fermentations were carried

out at 200 rpm, in the 3.7 liters Lab Fermenter (Bioengineering, Switzerland). The culture

volume was of 1.5 l, at the same initial glycerol concentration of 20g/l (Figures 9.1 and

Figures 9.3). Final PHB concentration was 1.116 g/l with a productivity of 0.0248 g/l*h at

30 °C, while the final concentration of PHB was 3.4 g/l with a productivity of 0.0771 g/l*h

at 33 °C. Therefore, the best temperature operation is the 33°C.

0

0.5

1

1.5

2

2.5

3

3.5

4

4.5

5

0 5 10 15 20 25 30 35 40 45 50

g/l

Time (hours)


Figure 9.2. Accumulation profile II of Bacillus megaterium cultivated in the 3.7 liters Lab


concentration of 50 g/l at 30°C and 200 rpm. Biomass, PHB.

Figure 9.3. Accumulation profile III of Bacillus megaterium cultivated in the 3.7 liters Lab


concentration of 20 g/l at 33°C and 200 rpm. Biomass, PHB. In the same form, glucose was used as only carbon source in order to compare the use of

glycerol as raw material. The initial glucose concentration used was of 20g/l at 33°C

(Figure 9.4). Final PHB concentration was 2.83 g/l with a productivity of 0.0783 g/l*h. A

comparison between obtained results for PHB production by mean of B. megaterium

using glucose and glycerol as substrates was made as shown in Table 9.2.

0

1

2

3

4

5

6

7

8

9

10

0 5 10 15 20 25 30 35 40

g/l

Time (hours)

0

1

2

3

4

5

6

0 5 10 15 20 25 30 35 40 45 50

g/l

Time (hours))


Figure 9.4. Accumulation profile IV of Bacillus megaterium cultivated in the 3.7 liters Lab

Fermenter (Bioengineering, Switzerland) with a culture volume of 1.5 l, initial glucose

concentration of 20 g/l at 33°C and 200 rpm. Biomass, Glucose, PHB

Table 9.2. Comparison between experimental results for glucose and glycerol

fermentation to PHB

Variable

GLUCOSE GLYCEROL

20g/l-T=33°C

20g/l-T=33°C

50g/l-T=30°C

20g/l-T=30°C

DCW(g/l)

5.08g/l 5.7g/l 7.8 g/l 4.7 g/l

PHB(g/l) 2.83 g/l 3.4g/l 2.356 g/l 1.116g/l

% Accumulation

55.7% 62.3% 30.2%.

23.74%.

Time of Max. PHB

production (hours)

36 42 32 45

Productivity

gPHB/l*h

0.0786 0.0771 0.07365 0.0248

Total substrate

consumption (g /l)

10.5 11.80 - 15.50

Yield P/S 0.30 0.29 - 0.07

The produced PHB was characterized. The peak got form the FT-IR spectrum (1728 cm-1) corroborates with the peak reported, indicated the presence of PHB using glycerol as

sole carbon resource.

0

5

10

15

20

25

0 5 10 15 20 25 30 35 40

g/l

Time (hours)


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B. Formation of polyesters by Pseudomonas oleovorans : effect of substrates on

formation and composition of poly-(R)-3-hydroxyalkanoates and poly-(R)-3-

hydroxyalkenoates. Appl Environ Microbiol 54 (1988), 2924±2932.

[24]. S.P. Valappil, S.K. Misra, A.R. Boccaccini, T. Keshavarz, C. Bucke, I. Roy. Large-

scale production and efficient recovery of PHB with desirable material properties, from the

newly characterized Bacillus cereus SPV. Journal of Biotechnology 132 (2007) 251–258

10. Conclusions

Commercially three qualities of glycerol were identified as the most important ones. Crude

glycerol with a purity ranging from 80-88 wt %, technical glycerol mainly found at 97 wt %,

and refined glycerol (USP or FCC grades) at 99.7 wt %. These three types of glycerol

differ significantly in the content of water, fatty acid residues, esters, and other organic

wastes. Also, some differences were found for the use of diverse feedstocks for biodiesel

production on the composition of the glycerol layer. Although, most of the first use oils

lead to not big differences in the glycerol layer, a completely different behavior was

observed for the glycerol obtained from WVO represented by low concentration of

glycerol and methanol with a high content of fats. On the other hand, based on the

traditional purification of glycerol, a flowsheet able to purify raw glycerol up to the three

commercial qualities above described was designed, simulated and economically

assessed. Results showed that not only quality requirements were successfully obtained

but also for the analyzed purification scale all the processes were profitable. Thus, a

homogenized raw material and a purification process were obtained in order to continue

the analysis of different possibilities of glycerol transformation to added-value products.

Acrolein, hydrogen, and 1,2-propanediol, are three of the most commercially important

products obtained from glycerol, due to their applications, established market, and sale

prices. Here the technological schemes to produce these compounds were designed,

simulated, and economically assessed. Thus, simulation results showed that all the

processes are technologically feasible reaching high purity of product. Also, acrolein

production was found to be viable at a purity of 92 wt %, but do not at a purity of 98.5 wt

%. Finally, both hydrogen and 1,2-propanediol production processes are also

economically viable, where the last one generates the highest profit margin.

In the past, important efforts have been made to introduce the biotechnological production

of 1,3-propanediol from glycerol to the industry. However, research tendencies were


focused on microorganism development and some process conditions analysis. The

drastic increment in the use of biodiesel caused an oversupply of glycerol in the market.

For this reason and in order to optimize its productivity, mass production of 1,3-

propanediol from glycerol needs additional analysis. Here, fermentation of glycerol with K.

pneumoniae was optimized using three different models and considering two

simultaneous goals: i) high volumetric productivity and ii) high 1,3-propanediol

concentration. In conclusion, the obtained results provide enough information to

understand the different possibilities for process intensification using this technology and

also to compare it with other new industrial alternatives for the utilization of glycerol as a

raw material.

Due to the low cost of raw glycerol, methanol recovery from glycerol implies low PCs.

Meanwhile, the three possibilities assessed for glycerol bioconversion showed that the

GPCs of fuel ethanol from raw glycerol are lower than the commercial price of fuel

ethanol. These facts show the potential for raw glycerol bioconversion to fuel ethanol

using E. coli. Also, the comparison carried out with a previous paper (which considers the

fuel ethanol production from sugarcane and corn in the Colombian case, [13]) shows that

the GPCs of fuel ethanol from raw glycerol can be as profitable as the production of fuel

ethanol from conventional raw materials as sugarcane. The latter is a completely

developed industry in Colombia.

Biodiesel and ethanol can be jointly produced using oil palm as sole source by mean of

processes integration, such as the biodiesel production with the ethanol production from

two feedstocks: lignocellulosic residues (empty fruit bunches and palm press fiber

produced during) and crude glycerol. Thus, alcohol is completely self-supplied by the

integrated process and low quantities of wastes are produced without any global

production of glycerol. Economical evaluation showed a higher biodiesel production cost

for the integrated process than the traditional biodiesel production process which uses

ethanol and palm oil as feedstocks. But the first one is a promising technology available to

build an autonomous biodiesel production plant with low waste levels. This process must

be economically improved by further analysis from a process design view point, based on

process simulation which showed be a powerful tool for performing processes integration.

10. Conclusions 197

Three technological schemes to produce PHB from crude glycerol were analyzed under

two fermentation conditions (i.e., using glycerol at 88 wt % and 98 wt %). In this work it

was found that it is better to use pure glycerol as feedstock for the production of PHB than

raw glycerol. This phenomenon is explained by the fact that the higher PHB yield reduces

the utility costs in the downstream process. The results shown here are important for the

industrial production of PHB using glycerol as a raw material. Currently, in the biodiesel

production the total profitability of any new project could be determined by the right use of

glycerol as a massive by-product. The proposed strategy to use pure glycerol as

substrate can be understood as a very interesting alternative since the final composition

in the glycerol streams depends on the source of the feedstock used for biodiesel

production. Also, most of the biotechnological alternatives to produce added value

compounds from glycerol are sensitive to contaminants in the raw material. Thus, several

technical and economical advantages as well as a more stable production of PHB are

obtained when a standardized raw material as pure glycerol is used.

Usage of raw glycerol, engineered Escherichia coli strains, and processes integration for

the production of optically pure D-lactic acid is an important alternative to transform the

by-produced glycerol during the biodiesel synthesis. Although five different configurations

for the fermentation stage were considered, in all cases the total production costs were

lower than its sale price. Thus, the whole process scheme for D-lactic acid production

could be considered as potentially profitable design. Also, it was found that the combined

effect of both high glycerol concentration and use of low quality glycerol in the

fermentation media, lead to the best economic performance. The results shown here are

important for the industrial production of D-lactic acid using glycerol as a raw material. On

the other hand, in the biodiesel production the total profitability of any new project could

be determined by the right use of glycerol as a massive raw material.

11. List of Publications and Submitted Papers

This chapter shows the published results throughout scientific meeting, papers, book

chapters, invited book chapters, and books. Also, a list containing the submitted papers

was made.

11.1 Published

Scientific Meetings

1. Posada JA, Higuita JC, Cardona CA. 2011. Optimal crude glycerol biorefinery

from biodiesel production to produce poly-3-hydroxybutyrate. World

Renewable Energy Congress 2011 – Sweden. Linköping University.

2. Posada JA, Quintero JA, Cardona CA. 2010. Energy and environmental

comparison among the production of fuel ethanol from crude glycerol, sugar

cane, and crop. IV International congress of science and technology of the biofuels. Bucaramanga, Colombia. November 30th - December 3rd

.

3. Posada JA, Cardona CA, Rincón LE. 2010. Sustainable biodiesel production

from palm using in situ produced glycerol and biomass for raw bioethanol. In:

Society for Industrial Microbiology. 32nd symposium on biotechnology for fuels

and chemicals. Clearwater Beach, Florida. Abril 19th-22nd.


Papers

4. Posada JA, Naranjo JM, López JA, Higuita JC, Cardona CA. 2011. Design

and Analysis of PHB Production Processes from Crude Glycerol. Process

Biochemistry. 46:310-317.

5. Posada JA, Cardona CA. 2010. Design and analysis of fuel ethanol

production from raw glycerol. Energy. 35(12):5286-5293.

6. Posada JA, Cardona CA. 2010. Análisis de la refinación de glicerina obtenida

como co-producto en la producción de biodiesel (Validation of glycerin refining

obtained as a by-Product of biodiesel production). Ingeniería y Universidad

14:2-27.

7. Posada JA, Orrego CE; Cardona CA. 2009. Biodiesel production:

Biotechnological approach. International Review of Chemical Engineering

(I.Re.Che.), 1(6):571-580.

Book Chapters

8. Posada JA, Cardona CA, Cetina DM, Orrego CE. 2009. Bioglicerol como

materia prima para la obtención de productos de valor agregado (Bioglycerol

as raw material to obtain added value products). En: CARDONA CA (ed).

Avances investigativos en la producción de Biocombustibles (Reasearching

advances for biofuels production). Manizales: Artes Graficas Tizan. p. 103-

127. ISBN: 978-958-44-5261-0

11. List of Publicatiions 201

Books

9. Cardona CA, Posada JA, Quintero JA. 2010. Aprovechamiento de

subproductos y residuos agroindustriales: Glicerina y Lignocelulósicos (Use of

agroindustrial wastes and by-products: Glycerin and Lignocellulosics).

Manizales: Artes Graficas Tizan. p. 218. ISBN: 978-958-44-7611-1

Invited Book Chapters

10. Posada JA, Rincón LE, Cardona CA. Integral Use of Palm Oil: Production of

Biodiesel and Added Value Compounds from Glycerin. In: Oil Palm:

Cultivation, Production and Dietary Components. Editor: Susan A. Penna.

Book Chapter Ed. Nova Publisher. Series: Agriculture Issues and Policies.

ISBN: 978-1-61122-201-2.

11.2 Submitted

Papers

1. Posada JA, Jaramillo JJ, Cardona CA. Glycerol fermentation to 1,3-

propanediol: Comparison among four culture configurations. Submitted to

Biochemical Engineering Journal.

2. Posada JA, Higuita JC, Pisarenko YA, Cardona CA. Design and economical

analysis of the technological scheme for 1,3-propanediol production from raw

glycerol. Submitted to Theoretical Foundations of Chemical Engineering.


3. Posada JA, Cardona CA. Comparison among three chemical processes for

glycerol conversion: Acroleine, Hydrogen and 1,2-Propanediol. Submitted to

Bioresource Technology.

4. Posada JA, Rincón LE, Cardona CA. Sustainable biodiesel production from oil

palm: crude glycerol and biomass conversion to raw-bioethanol. Submitted to

Applied Biochemistry and Biotechnology.

5. Posada JA, Quintero JA, Cardona CA. Comparación económica y ambiental

entre la producción de etanol a partir caña de azúcar, maíz y glicerol crudo.

Submitted to Revista de Ingeniería Química-UIS.

6. Posada JA, Quintero JA, Cardona CA. Comparison among three technologies

for biodiesel production from Jatropha seeds. Under review by the advisor.

7. Posada JA, Cardona CA. Possibilities of glycerol conversion as a sole raw

material: a review. Under review by the advisor.

8. Posada JA, Cardona CA, Gonzalez R. Design, simulation, and economic

assessment of D-lactic acid production process from raw glycerol using

engineered Escherichia coli strains. Under review by the advisor.

9. Posada JA, Naranjo JM, López JA, Higuita JC, Cardona CA. Poly(3-

hydroxybutyrate) production by Bacillus megaterium using glycerol as

substrate: experimentation and process simulation. Under review by the

advisor.

10. Posada JA, Cardona CA. Design, simulation, and economic assessment of

succinic acid production process from raw glycerol. Under review by the

advisor.

11. List of Publicatiions 203

11. Posada JA, Cardona CA. Design, simulation, and economic assessment of

propionic acid production process from raw glycerol. Under review by the

advisor.

Documents

ANALISIS COMPARATIVO DE USOS DE GLICEROL