Recovery of bio-based butanol
Recovery of
bio-based butanol
You are cordially invited to
the public defence of
my doctoral thesis
Recovery of
at 10:00h,
Monday, June 11th, 2012
in the Senaatszaal in
the Aula of the Delft
University of Technology,
Mekelweg 5, Delft.
A short presentation will
precede the defence at 9:30h.
Following the defence
a lunch is given at the
botanical garden behind
the Kluyver Laboratory,
Julianalaan 67 Delft.
Arjan Oudshoorn
Oudshoorn_Omslag.indd 1
Arjan Oudshoorn
Arjan Oudshoorn
14-05-12 15:25
Recovery of bio-based butanol
Arjan Oudshoorn
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Author: A. Oudshoorn ©
Recovery of Bio-Based Butanol
ter verkrijging van de graad van doctor
aan de Technische Universiteit Delft;
op gezag van de Rector Magnificus prof. ir. K.C.A.M. Luyben,
voorzitter van het College voor Promoties
in het openbaar te verdedigen op maandag 11 juni 2012 om 10:00 uur
Arjan Oudshoorn
Master of Science in Biochemical Engineering
geboren te Leiden
Dit proefschrift is goedgekeurd door de promotor:
Prof. dr. ir. L.A.M. van der Wielen
Dr. ir. A.J.J. Straathof
Samenstelling promotiecommissie:
Rector Magnificus
Prof. dr. ir. L.A.M. van der Wielen Technische Universiteit Delft, promotor
Dr. ir. A.J.J. Straathof
Technische Universiteit Delft, copromotor
Prof. dr. G. Eggink
Wageningen Universiteit
Prof. dr. F. Kapteijn
Technische Universiteit Delft
Prof. dr. ir. M.C. Kroon
Technische Universiteit Eindhoven
Dr. J.-P. Lange
Shell Global Solutions International B.V.
Prof. dr. ir. G.J. Witkamp
Technische Universiteit Delft
Dit onderzoek werd financieel ondersteund door het ministerie van economische zaken en
de B-Basic partner organisaties ( via B-Basic, een publiek-privaat NWOACTS programma (ACTS = Advanced Chemical Technologies for Sustainability).
ISBN 978-90-5335-561-9
Table of Contents
Chapter 1
Introduction to bio-based butanol recovery
Chapter 2
Assessment of options for selective 1-butanol recovery
from aqueous solution
Chapter 3
Exploring the potential of recovering 1-butanol
from aqueous solutions by liquid demixing
upon the addition of carbohydrates or salts
Chapter 4
Adsorption equilibria of bio-based butanol solutions
using zeolite
Chapter 5
Desorption of butanol from zeolite material
Chapter 6
Short-cut calculations for integrated product
recovery options in fermentative production
of bio-bulk chemicals
Chapter 7
Outlook to bio-based butanol recovery
Product recovery is crucial for fermentative butanol production. Bio-based fermentative
butanol production systems benefit from answering the question to what extent, and in
which form, integration between fermentation and product recovery should be applied. The
product recovery is applied in order to minimize the impact of butanol inhibiting during
fermentation. This thesis addresses the product recovery methods applicable to butanol
recovery. Two techniques are investigated in more detail, namely liquid demixing basedrecovery and adsorptive-recovery. This work further provides a basis for process
optimization for integrated product recovery and finally determines the economic potential
of butanol production systems.
Butanol can be removed from the fermentation broth by means of direct phase transition
(e.g. distillation or freeze crystallization), or by using auxiliary phase (e.g. extractive
recovery or adsorption). The assessment of butanol recovery from aqueous solutions,
shown in chapter 2, illustrates the wide range of recovery techniques which can be applied.
Key efficiency aspect for product recovery is the selectivity of the recovery technique. The
selection of the correct product recovery technique is paramount as is shown by the range in
energy requirements for butanol recovery, ranging from 1.3 MJ/kg to 145 MJ/kg.
Selectivity can be used as a parameter to qualify the energy demand of a production system.
Selectivities of recovery can be a function of the initial product concentration in the
fermentation broth, but to illustrate, for butanol recovery the highest product recovery
selectivities are reported for extraction (S = 4100), liquid demixing (S = 750) and
adsorption (S = 630). For the commonly applied distillation selectivity is lower, (S = 72).
The selectivity is based on the liquid-gas phase equilibria. This lower selectivity implies
that distillation will be carried out as a multistage operation and also that heat integration
should be applied.
This thesis does not investigate the recovery by organic solvent extraction further as the
method is extensively reported in literature and the extractive phase capacity for butanol is
extremely limited, <0.003 kg butanol per kg organic extractive phase.
Liquid demixing is explored in chapter 3 as recovery method for butanol. The recovery
makes use of the distinctive phase behavior of butanol-water mixtures, which shows a two
liquid phases equilibria for a wide range of compositions. The organic phase of these two
can be seen as a high concentrated product phase. Our analysis shows the liquid equilibria
to be susceptible to chemical species used in fermentative production systems, namely salts
(LiCl, NaCl and CaCl2) and carbohydrates (glucose, fructose and sucrose). The effects of
carbohydrates on the liquid phase equilibria are less pronounced than the effect of the salts.
The amount of salts needed for product recovery exceeds 250 g kg-1 and are too high for
practical consideration. The amount of salt needed for product recovery goes down
significantly the moment the butanol concentrations in the fermentation are closer to
saturated conditions. An increase in solvent tolerance of microbial species are a favorable
development for liquid-liquid based recovery, although butanol fermentations will not come
close to saturated product concentrations 74 g L-1.
Adsorptive recovery by high silica zeolites shows the zeolite affinity for butanol to be
dependant on the zeolite structure and on aluminum content. ZSM-5 high silica zeolite,
with a narrow pore structure, shows a very steep adsorption isotherm, indicating the
beneficial effect of limited pore volume and hydrophobic nature of the pores for the
selectivity of recovery. The butanol adsorption capacity for CBV28014 is actually limited
by the available pore volume. Multi-component adsorption behavior for the investigated
zeolites, when dealing with Acetone-Butanol-Ethanol (ABE) fermentations, can be
estimated from the pure component adsorption isotherm in water by either simple multicomponent Langmuir model or ideal adsorbed solution model. When modeling adsorption
from fermentation broths, the acid content, e.g. butyric acid, needs to be taken into account,
and should be present as an additional chemical species in the extended adsorption model.
Desorption of butanol from zeolite material by means of temperature operation is possible
and is shown in detail in chapter 5. Differential Scanning Calorimetry (DSC) and Thermo
Gravimetric Adsorption (TGA) experiments show the heat effect of desorption of butanol
from CBV901 and CBV28014 to be slightly above the enthalpy of evaporation, namely
1080 and 1160 J.g-1. Also the structurally more hydrofobic CBV28014 shows the least
amount of water adsorption. The confined nature of the CBV28014 structure does show
significantly slower desorption rates, with overall mass transfer coefficient being up to 10
times slower, compared to butanol desorption from the wider pore CBV901. The low heat
capacity of silica material of around 1 J.g-1, the adsorptive phase capacity for butanol of
approximately 0.1 g.g-1, the high product recovery selectivity and thermal stability of the
material make adsorption based recovery a viable method for butanol recovery.
Carbon dioxide has an effect on the adsorption of butanol on high silica material in both
liquid as well as in gas phase systems. Carbon dioxide can thus be used as a displacement
agent to allow adsorptive phase regeneration. However, just as for most organic phases
applied in extractive based recovery, the butanol content in carbon dioxide is extremely
The process evaluation carried out in chapter 6 show that for butanol both fermentation and
product recovery are in the same range of importance. In-situ or integrated product
recovery is necessary in order to optimize the expenditure of both unit operations. The
capital expenditure (capex) and the operational expenditure (opex) for butanol recovery
systems allow estimation of the costs related to bio-based butanol and these costs are
currently estimated to be between 1.5 and 2 euro per kg.
Terugwinning van bio-butanol
Productisolatie is belangrijk voor fermentatieve butanolproductie. Bio-gebaseerde
fermentatieve butanolproductieprocessen halen voordeel wanneer antwoord wordt gegeven
op de vraag hoeverre en in welke vorm integratie tussen de fermentatie en de
productwinning moet plaatsvinden in het productieproces. Productisolatie wordt toegepast
om de inhibiterende werking van butanol op de fermentatie te minimaliseren. Dit
proefschrift behandelt productisolatiemethodes die toegepast kunnen worden op isolatie van
vloeistofontmenging en adsorptiegebaseerde productisolatie. Dit proefschrift beschrijft ook
een basis voor optimalisatiemethodologie voor geïntegreerde product processen voor
butanolproductiesystemen zien.
Butanol kan verwijderd worden uit een fermentatiemedium door middel van directe
faseverandering (bijv. destillatie of vrieskristallisatie), of door gebruik te maken van een
isolatiemogelijkheden uit waterige oplossingen, gepresenteerd in hoofdstuk 2, illustreert het
brede scala dat kan worden toegepast. Vanuit het oogpunt van efficiëntie is de selectiviteit
die gehaald wordt in de gebruikte methode cruciaal. De selectie van een correcte productisolatietechniek is van het hoogste belang, zoals wordt geïllustreerd door de grote variatie
in de energiebehoeftes van de verschillende technieken die voor butanolwinning kunnen
worden gebruikt, variërend van 1,3 MJ/kg tot 145 MJ/kg. Selectiviteit kan als parameter
gebruikt worden om een uitspraak te doen over de energiebehoeften van een
productieproces. De selectiviteit van de isolatie van het product kan als functie van de
initiële concentratie in de fermentatie uitgedrukt worden. Ter illustratie, de hoogste product
selectiviteit (S) die is gerapporteerd is voor extractie S = 4100, voor vloeistof ontmenging S
= 750 en voor adsorptie S = 630. Voor de standaard toegepaste destillatie is de selectiviteit
lager, S = 72. Deze selectiviteit komt voort uit het vloeistof-gas fase-evenwicht. Deze
lagere selectiviteit impliceert dat destillatie alleen uitgevoerd kan worden als meertraps
operatie en dat warmte-integratie moet worden toegepast.
Dit proefschrift beschrijft isolatie van product door extractie met organische oplosmiddelen
niet in detail, omdat extractie uitvoerig gerapporteerd wordt in de literatuur. Daarbuiten valt
nog op te merken dat de capaciteit voor butanol van de extractieve fase is extreem beperkt,
<0,003 kg butanol per kg organische fase. Vloeistofontmenging is onderzocht in hoofdstuk
3 als mogelijke isolatiemethode voor butanol. De terugwinning maakt gebruikt van het
onderscheidende fasegedrag van mengsels van water en butanol, wat vloeistofontmenging
laat zien over een breed concentratiegebied. De organische fase kan hier gezien worden als
een geconcentreerde productfase. De analyse laat zien dat de vloeistofevenwichten
beïnvloed worden door chemicaliën die ook tijdens de fermentatie worden gebruikt,
namelijk zouten (NaCl en CaCl2) en koolhydraten (glucose, fructose en sucrose). Het effect
van koolhydraten op de vloeistofevenwichten is minder uitgesproken dan het effect dat
zouten hebben op de evenwichten. De hoeveelheid zout die nodig is voor fasescheiding is
meer dan 250 g per kg en is vanuit praktisch oogpunt te hoog. De hoeveelheid benodigd
zout gaat significant omlaag op het moment dat de butanolconcentratie in het
fermentatiemedium dicht in de buurt komt van verzadiging. Een verbetering in tolerantie
van het micro-organisme voor organische oplosmiddelen is een gunstige ontwikkeling op
vloeistof-ontmenging gebaseerde scheidingstechnieken, al zal butanolfermentatie niet
makkelijk in de buurt kunnen komen van de verzadigingsconcentratie, 74 g L-1.
Adsorptie met behulp van silica zeoliet laat zien dat de affiniteit voor butanol van de zeoliet
afhangt van de zeoliet structuur en de aanwezige hoeveelheid aluminium. ZSM-5 hoog
silica zeoliet, met nauwe porie-structuur, laat een erg scherpe adsorptie-isotherm zien, wat
het voordeel van beperkte porievolume en hydrofobe aspecten van de poriën op de
selectiviteit van de scheiding illustreert. De adsorptiecapaciteit van CBV28014 voor
butanol is gelimiteerd door het beschikbare porievolume. Multi-component adsorptiegedrag
door de bestudeerde zeolieten van Aceton-Butanol-Ethanol (ABE) mengsels laat zien dat de
adsorptie kan worden voorspeld aan de hand van de adsorptie-isothermen van de pure
componenten indien gebruik wordt gemaakt van een simpel multi-component Langmuirmodel of het ideal adsorbed solution (IAS) model. Voor het modelleren van de adsorptie uit
fermenatiemedium is het belangrijk de aanwezige zuren in beschouwing te nemen, bijv.
boterzuur, en deze moeten als extra component aan het model worden toegevoegd.
Desorptie van butanol uit zeoliet materiaal door middel van temperatuurverandering is
mogelijk en wordt in detail behandeld in hoofdstuk 5. Differential Scanning Calorimetry
(DSC) en Thermo Gravimetric Adsorption (TGA) experimenten met CBV901 en
CBV28014 laten zien dat het warmte-effect van desorptie van butanol respectievelijk 1080
and 1160 J.g-1 bedraagt, iets boven de verdampingsenthalpie van butanol. Ook laat het
meer hydrofobe CBV28014 significant minder wateradsorptie zien. De structuur van de
CBV28014 veroorzaakt meer massatransportlimitatie en de massatransportcoëfficiënt voor
de totale overdracht is tot 10 keer lager dan voor desorptie van butanol uit het meer open
CBV901. De lage warmtecapaciteit van silicamateriaal van ongeveer 1 J.g-1, de adsorptie
capaciteit voor butanol van ongeveer 0.1 g.g-1, de hoge scheidingsselectiviteit en de
thermische stabiliteit van het adsorptiemateriaal maken adsorptiegebaseerde processen een
reëel toepasbare methode voor butanolproduct-isolatie.
Koolstofdioxide heeft een invloed op de adsorptie van butanol door silica materiaal in
zowel vloeistof- als gassystemen. Koolstofdioxide kan zodoende gebruikt worden als
verdringingsmiddel om de adsorptiefase te regenereren. Echter, net zoals voor de meeste
organische oplosmiddelen die gebruikt worden in extractieve product-isolatieprocessen, is
de oplosbaarheid voor butanol in koolstofdioxide beperkend.
Een procesevaluatie is uitgevoerd in hoofdstuk 6. Deze laat zien dat voor butanol zowel de
fermentatie als de product-isolatie qua ordegrootte een gelijke invloed hebben op de
proceskosten. In-situ of geïntegreerde product-isolatie is nodig om de kosten van beide
operaties te kunnen minimaliseren. De investeringskosten en de operationele kosten voor
productieprocessen voor uit fermentatie verkregen butanol geven een uiteindelijke
kostenraming die ligt tussen de 1,5 en 2 euro per kg butanol.
Chapter 1: Introduction to bio-based butanol recovery
1.1. Bio-based microbial chemicals and fuels production
The world production of chemicals and fuels is at this time predominantly based on the
conversion of non renewable raw materials, such as coal, natural gas and oil. Inevitably
these resources will be depleted. The timescale on which these resources are depleted are a
matter of debate, but our society will run out of these natural resources eventually.
From the standpoint of depletion, switching to a renewable resource for the production of
chemicals and fuels is a logical step. Renewable resources are almost all derived from
sunlight. Sunlight, wind, geo-thermal or hydro-electric systems can be used for energy
production. The production of most chemicals and transportation fuels do require carbon
based resources. The abundant carbon based material on the planet besides carbon dioxide
is biomass. Plant material, algae and microbial biomass all can be converted into chemicals
and fuels. Conversion of renewable carbon based feedstocks can take place by a wide range
of processes from thermo-chemical and catalytic chemical conversion to enzymatic and
microbial conversion.
Bio-based microbial production of bulk chemicals has existed since the start of the 20th
century. Due to the rise and successfulness of the petrochemical industry the competition
was lost after the 1950’s and bio-based bulk production of most chemicals ceased.
Currently, conversion steps are again increasing in competitiveness. Successful bio-based
processes are on the market, with 1,3-propanediol, citric acid, lactic acid and ethanol being
prime examples. 1-Butanol is currently joining these examples
1.2. Microbial 1-butanol production
Historically acetone-butanol and ethanol (ABE) fermentation was already occurring on
large scale up to the late 1950’s. Butanol as a biofuel has favorable properties over the
already existing ethanol production in its higher combustion energy on mass basis and its
lower polarity, allowing blending in biodiesels. Further, butanol is scientifically an
interesting component with its hydrophobic carbon chain and its hydrophilic alcohol group,
its interesting phase behavior with water, and its low saturated vapor pressure.
Microbial butanol fermentation occurs in water and the recovery of low concentration
butanol from an aqueous phase is non-trivial and requires the investment of equipment and
energy. In contrast industrial petrochemical conversions usually occur in an organic or
vapor phase and its separation of choice, distillation, is predominantly the separation of
organic products from other organic components, while water is mostly used in the form of
steam as heat transport agent. With the switch towards aqueous fermentative production,
the product separation conditions significantly change. What is currently an optimal
recovery need not be the optimum recovery for future renewable processes.
1.3. Towards bio-based butanol production
Large scale fermentations can be severely hampered by product inhibition. This means the
microbial production is being limited as the produced products themselves negatively
influence the microbes. Product toxicity effects can be limited if the product is continuously
removed from the production system. Product removal can come in different process
configurations and can be applied in-situ or in a multi-step integrated process. All of this
applies to microbial butanol production, and as the transition to a more renewable resource
based society is still in an early stage, it is the right time to investigate butanol recovery
techniques, besides the currently dominant technique of distillation. This assessment can
than allow a subsequent determination of the potential of bio-based butanol production,
given the wide range of integrated production options.
This thesis thus systemically investigates the recovery of fermentative produced butanol
from aqueous solutions; further investigates some specific promising techniques; provides a
basis for integrated process optimization; and finally provides a basis for the determination
of the economic potential of butanol production.
The approach used to investigate the microbial production of butanol can be used as a
roadmap for the investigation of other microbial produced chemicals.
1.4. Outline of this thesis
In chapter 2 an assessment of selective separation techniques of butanol is made using a
systematic approach, starting at the phase behavior of butanol-water mixtures. This
assessment shows the current state of the recovery techniques and provides the background
for the later chapters.
In chapter 3 the liquid-liquid demixing based recovery of butanol is explored, as it is
identified as a currently underexplored technique in chapter 2.
In chapter 4 adsorptive recovery of butanol on high silica material is investigated, as
adsorption is identified as a selective recovery technique in chapter 2.
As follow up on chapter 4, chapter 5 shows the thermal desorption characteristics of
butanol from high silica material. Desorption of products from their sorbent is less
frequently studied than adsorption, while its importance in relation to the overall recovery
process is very high.
Chapter 6 shows the integrated production aspects of butanol production and provides the
breakdown of capital and operational expenditure of its fermentative production. The actual
butanol product concentration in the fermentation and downstream processing plays a
crucial role in process optimization. This method is also applied to lactic acid and phenol
Chapter 7 provides an outlook on the product recovery of microbial produced butanol and
microbial production of bulk chemicals in general.
Chapter 2: Assessment of options for selective 1-butanol recovery from aqueous
The microbial production of 1-butanol occurs in aqueous fermentation broth, with up to ~
20 g/L of product. Efficient recovery of butanol from this dilute aqueous phase determines,
to a large extent, the efficiency of the production process. Starting from the thermodynamic
(phase) properties of butanol and water systems, this paper presents a structured approach
to determine the key characteristics of various butanol recovery methods. Analysis of
reported separations, combined with fundamental phase properties, has resulted in both the
characterization of the selectivity of recovery and estimations of the energy requirement
during product recovery for a variety of recovery methods. Energy-efficient systems for the
recovery of butanol from aqueous solution are pervaporation- and adsorption-based
techniques. The applied method predicts the recovery energy requirement for both
techniques to be < 4 MJ/kg of butanol, which, on an energy basis, is similar to ~ 10% of the
internal combustion energy of butanol.
Published as: Assessment of Options for Selective 1-Butanol Recovery from Aqueous
Solution, Arjan Oudshoorn, Luuk A.M. van der Wielen, and Adrie J.J. Straathof, Ind. Eng.
Chem. Res. 2009, 48, 7325–7336 7325
2.1. Introduction
The anaerobic production of organic solvents by biological conversion of renewable
feedstocks dates back to 1861 to Pasteur. Glycerol fermentations with its main products
butyrate and butanol were later described by A. Fitz in 1876. Acetone, 1-butanol, and
ethanol (ABE) fermentation was second only to ethanol production as a biological
production route in the beginning of the 20th century. In the mid-20th century, economic
factors and new petrochemical production methods led to the decline of the ABE
fermentation industry, although in South Africa, the Soviet Union, and China,1-3 production
continued beyond this point. Increasing interest in sustainable industry, the increase in oil
prices and renewable feedstock utilization has led to renewed attention for ABE
fermentation from industry4 and academia,5 particularly for 1-butanol (hereafter referred to
as butanol), although new developments also include 2-butanol and iso-butanol. Butanol
can be used as a solvent, as a precursor for chemical synthesis, or as a biofuel. From a
biofuel perspective, butanol has some advantages over ethanol. Butanol has a 31% higher
combustion value, compared to ethanol. Also, the chemical properties of butanol, such as
chain length, lower volatility, and polarity, allow blending in biofuels6 more readily than
ethanol does.
ABE fermentation is performed by a large variety of Clostridia strains at 25-37 oC and 1
atm. In ABE fermentations, butanol is usually the main product.1,7 The final ABE
composition can vary, but the maximum achievable total solvent concentration is ~ 20 g/L,
mostly because of severe product inhibition and toxicity by butanol. Process synthesis
approaches have focused on separation of the ABE mixture, without involving the
fermentation.8,9 We suppose that metabolic engineering approaches will be increasingly
successful in minimizing acetone and ethanol formation in the future.
However, the achievable butanol concentration will remain modest. Therefore, we will
focus here on in situ recovery of butanol without considering acetone and ethanol. By
removing product during the fermentation, the productivity per volume of fermenter and
per amount of cell mass can increase significantly. This concept has been commercially
applied for ethanol production using the Biostil process, and other options (such as
extractive lactate fermentation) are awaiting implementation. 10
The fermentative production and recovery of butanol can be performed according to
various schemes, allowing direct and indirect cell contact and internal and external product
recovery.11 The scheme shown in Figure 1 gives the best control possibilities and, therefore,
is assumed to be the most suitable option for a large-scale continuous process. The product
capture step shown in Figure 1 consists of a downstream operation that leads to a smaller,
more-concentrated butanol stream and a large, more dilute aqueous stream for recycle to
the fermentation. This aqueous stream may contain the microbial cells; however,
preferably, these are retained in the fermenter. This issue is not discussed in this paper,
which will focus on the capture step. Complete recovery of butanol in the capture step is
not necessary, assuming that the remaining aqueous stream, except for a small purge
fraction, can be recycled to the fermentation.
The integrated butanol production system can be subunit of a (bio)refinery, where many
streams will be available for further process integration. To avoid bias in evaluating the
actual performance of the integrated system, stream integration with streams that do not
stem from the butanol fermentation is not applied.
The separation of butanol and water in this operation will be based either on differences
between the pure component physical properties or on their inherently different interaction
with a third chemical species or auxiliary material. A favorably chosen ternary species, or
mass-separating agent, will form an auxiliary phase in the capture step and will facilitate
the recovery. To achieve product specifications, a final purification step may be required.
Figure 1. Butanol capture operation in its process context.
Ideally, the main costs in a butanol production process are due to feedstock consumption.
Using cell retention and optimized fermentation techniques, the butanol yield on sugars can
be maximized and feedstock costs will be fixed. Other large cost factors are due to the
fermentation equipment, the capture equipment, and the overall energy consumption during
recovery. An efficient capture operation will minimize all these costs, including the
fermenter costs, if a less-inhibiting butanol concentration can be maintained in the
fermenter, according to the recovery scheme of Figure 1. Therefore, it will be crucial to
select the type of operation for the capture step.
Distillation is the traditional recovery option, but the literature covers a wide range of
alternative recovery options. These usually allow no straightforward comparison, because
feed conditions and underlying assumptions vary from case to case.
Recovery information on an isolated separation step also provides limited information
about its suitability within a process. Performing complete process designs on all possible
systems is time-consuming. Ranking the possible butanol recovery options, however, is
necessary for a rational choice. As mentioned previously, the energy consumption is a key
cost factor of the recovery operation. We will use it as a parameter to rank recovery
alternatives. The energy consumption influences not only operational costs but also
investment costs through heat-exchange area. Therefore, our strategy to allow quantitative
comparison of recovery alternatives is 3-fold and takes the mass and energy flows during
the recovery operation into account separately. First, in Section 2, a phase-transition-based
framework is used to compile and structure butanol and water separation possibilities on a
thermodynamic basis. Section 3 characterizes the recovery options by their performance in
separating butanol from water, using selectivity as a parameter. In Section 4, a short-cut
method is presented, to allow ranking of the recovery options on an energy basis without
having to perform labor-intensive process designs, while keeping the amount of necessary
property data low.
Table 1. Recovery Operations for Butanol from Aqueous Solution
new phase
P/T shift
distillation 13,14
P/T shift
liquid demixing
P/T shift
P/T shift or
composition change
P/T shift
composition change
gas stripping 1,20-23
composition change
composition change
liquid demixing 28
perstraction 25,29
composition change
adsorption 30-34
Table 2. Phase-Transition Properties of Pure Butanol and Watera
Tm [oC]
Tb [oC]
Data taken from ref 36.
2.2. Thermodynamics of Butanol and Water Mixtures
The recovery of butanol from aqueous solution is governed by the phase behavior of
butanol during separation. After separation, the butanol product phase can be a vapor,
liquid, solid, or supercritical phase. The separation operation can be classified as an energy,
mass, or kinetic separation,12 or a combination thereof. In energy-based separations,
temperature and pressure are used as operational variables to introduce energy into the
system. In mass-based separations, an auxiliary phase is introduced to facilitate separation.
Here, variations in the chemical composition -and, to some extent, temperature and
pressure- drive the separation. Differences in the transport properties of the components
allow kinetically based separations. Membrane-aided separations are clear examples of
kinetic separations.
First, an overview of literature on butanol recovery operations from (cell-free) fermentation
broth and model solutions is shown in Table 1. The recovery options have been primarily
arranged by operating parameter. Second, a further subdivision is made by means of the
butanol product phase characteristics. For P/T driven systems, membrane operations also
are distinguished.
2.2.1. Pure Components. Liquid butanol and water can undergo phase transitions to vapor
or solid phases. The enthalpy of these transitions is shown in Table 2. On a molar basis, the
values for butanol are slightly higher; however, when recalculated on a mass basis, the
enthalpies for water are ~3 times higher than that for butanol. The pure-component
saturated vapor pressures, as a function of temperature for butanol and water,35 are shown
in Figure 2. Water has a 2-4-fold higher saturated vapor pressure.
2.2.2. Binary Systems. According to Gibbs’ phase rule, the number of independent
variables required to identify the intensive state of the system is 4 - the number of phases
present for binary systems of butanol and water. This means only two degrees of freedom
are available for biphasic systems, and the independent variables (temperature and
pressure) can be used to define the composition of the system. To avoid three dimensional
diagrams that cannot be read with a high degree of accuracy, only two-dimensional
diagrams are shown here. The temperature versus composition diagram for butanol and
water systems at 1 bar is shown in Figure 3. The vapor-liquid, liquid-liquid, and solidliquid equilibrium lines are based on original data points.37-39 By shifting the pressure, all
individual equilibrium lines shift. As previously mentioned, ABE fermentations are usually
performed in a temperature range of 25-37 oC. Starting with a butanol concentration of ~ 20
g/L, or 0.5 mol %, moving either up or down in temperature will lead primarily to a phase
transition of water, rather than a phase transition of butanol. Unfortunately, this would
recover ice rather than butanol crystals upon freezing. Upon boiling, the vapor is somewhat
enriched in butanol.
In Figure 3, the temperature effect on the mutual solubility of butanol and water can also be
seen. The aqueous solubility of butanol ranges from 2.8 mol % at 0 oC to 1.5 mol % at 60
Figure 2. Saturated vapor pressure of butanol and water. (Data taken from ref 35).
The mutual solubilities of butanol and water increase with an increase in pressure (see
Figure 4).40 The temperature effect is shown by the multiple P-x curves. The pressures
required to influence the solubility effectively are beyond those applied in conventional
industrial large-scale systems. Conceptually, high-capacity, small-volume devices could
allow pressure as an operational parameter for separation. The atmospheric vapor-liquid
equilibrium in Figure 3 shows an azeotropic point at x ≈ 0.25. Azeotropic behavior can
severely complicate the direct distillation of a mixture. The liquid-liquid-vapor (L-L-V)
equilibrium curve of a binary mixture is monovariant for a given composition when the
temperature is fixed. The equilibrium vapor pressure can be determined for any
composition and is shown in Figure 5. The pure-component vapor pressures of water and
butanol are given at x= 0 and x = 1, respectively.
2.2.3. Ternary Systems. Besides temperature and pressure, an auxiliary component can be
exploited to drive the separation of butanol from water. This third chemical species may
lead to additional phase equilibria. Also, systems that contain more than three chemical
species can be created; however, this investigation will be limited to ternary systems.
Introduction of a fourth (or even more) species will require more-complex regeneration
procedures and involve more phase equilibria. To illustrate the phase behavior of butanol
and water mixtures, four types of ternary phase diagrams are shown subsequently. The
ternary mixtures involve a solid, liquid, or gaseous compound with various mutual
solubilities. Because the concentration of the butanol in the aqueous phase is close to the
outer edge of the ternary diagram, the examples have arbitrary axis units, to enhance clarity
in illustrating the various types of phase behavior. The bottom axis of the diagrams shows
the water-butanol binary mixture. Three points on this binary axis are common to all
diagrams. The point on the left-hand side shows the liquid butanol concentration at feed
composition, xfBuOH. The two remaining points give the liquid-liquid (L-L) equilibrium
compositions. The addition of a ternary species to the initial butanol-water mixture will
change the overall composition from point xfBuOH in a straight line to the top of the triangle.
In each case, in the ternary phase diagrams, a tieline can be reached, such that the mixtures
split into a relatively butanol-poor aqueous L1 phase and a relatively butanol-rich second
phase. The ternary species shown in the diagrams are not necessarily the most efficient
separating agent, but they are common in literature and representative for other ternary
species. Ternary Systems with a Soluble Solid Species.
Figure 6 shows the butanol-water-potassium iodide (KI) system. KI is an example of a solid
that can dissolve partially in the aqueous phase and partially in the organic phase. The two
solubility points of the salt are present on the outer edges of the diagram. The mutual
solubility of butanol and water changes with the third species such that salting-out can
A similar diagram can be expected with other inorganic compounds like NaCl or polar
organic compounds such as amino acids.41,42 Figures 7 and 8 show experimental solubility
data for butanol and water solutions with NaCl and KI. The aqueous butanol equilibrium
concentration decreases significantly when the salt content increases. Simultaneously the
organic phase becomes less attractive for water. Predictive models for the salt effect on
solubility have been developed.43,44 Besides inorganic solids, organic solids can also have
an effect on 1-butanol and water mixtures.41 Ternary Systems with a Gaseous Species. A gaseous species can also be present
as a separate (gas) phase (see Figure 9). The gaseous phase will contain both butanol and
water. A common stripping gas for microbiological systems is nitrogen. Nitrogen will
dissolve in both the aqueous and organic phase. The effect on the L1 and L2 composition is
small and the butanol and water contents in the gaseous phase are low. Besides allowing
butanol to transfer to the gas phase under ambient conditions, no direct benefit, such as
enrichment of the gas phase, relative to the vapor phase in the binary system, is expected. A
gas that is inert, inexpensive, and insoluble in the aqueous solution is preferred. The Henry
coefficient can be used to describe the equilibrium partitioning of the compounds between
the gas and liquid phases. Ternary Systems with a Water-Immiscible Liquid Species. The ternary phase
behavior of a water-immiscible liquid species with complete miscibility of butanol and the
ternary component is shown in Figure 10, using octanol as the example. Two solubilities on
the water-octanol axis are indicated. The organic phase (L2) is present at any binary
butanol-octanol composition. Upon adding octanol to the feed composition xf BuOH a liquidliquid phase split occurs with a relative high butanol: water ratio in the organic phase.
Figure 3. T-x-y data for the binary mixture of butanol and water at 1 bar. Experimental data
points are indicated by markers ((▲) vapor-liquid
equilibrium,38 (♦) liquid-liquid39 equilibrium, and (●) solid liquid equilibrium).38 Lines
are included to guide the eye. The square markers at
the outer edges show the pure component fusion and boiling points for water (x = 0) and
butanol (x = 1).
Figure 4. Butanol and water solubility isotherms. (Data taken from ref 40.) The
temperature (displayed in degrees Celsius) is indicated at each isotherm.
The ternary phase data usually available for liquid ternary systems are expressed by
partition coefficients. The partition coefficients of butanol and water between organic and
aqueous phase are shown in Table 3 for various liquid nonpolar ternary species. The
partition coefficient morg/aq
BuOH is the equilibrium ratio between the mass fraction of butanol in
the organic phase and the mass fraction in the aqueous phase. Any m org/aq
BuOH value of > 1
indicates a preferable partitioning of butanol toward the organic phase. Table 3 gives some
interesting candidates. The solubility of water in the solvent also can be expressed as a
partition coefficient, m org/aq
H 2O . This coefficient is needed to determine the overall selectivity
of the separation. Selectivity can be defined as the ratio of the two partition coefficients. In
integrated systems, such as that shown in Figure 1, the aqueous raffinate is recycled to the
fermentation section and therefore the solubility of the organic solvents in the aqueous
phase, sSolvent
, must be considered. The solubility should be low, because it can disturb the
fermentation, if the solvent is not recovered from the aqueous raffinate. Of course, the
solvent should be easily recoverable from the extract to regenerate it. Some organic
solvents do not mix completely with butanol. In these situations, the phase diagram may
resemble Figure 9 rather than Figure 10, and the butanol solubility in the ternary species,
s org
BuOH , dictates the capacity for butanol and is an important parameter that influences the
total solvent requirement. Another situation occurs, for example, for the ternary mixture of
the ionic liquid [C4mim][NTf2] with butanol and water at 15 oC and atmospheric pressure.
This shows demixing in all three individual binary mixtures; however, complete mixing of
water, butanol, and the ionic liquid occurs over a large composition range.46 We refer to the
literature for the phase diagram.
Figure 5. (♦) Saturated total vapor pressure and partial vapor pressure of (□) butanol and
(Δ) water at T = 50 oC in binary mixtures. (Data taken from ref 35.) In the mole fraction
range without data, liquid demixing occurs.
Figure 6. Schematic ternary phase diagram of butanol and water with a soluble (solid)
species (example: potassium iodide (KI)).
Figure 7. Salt effect on butanol solubility in the aqueous phase, as a function of salt content
at T = 25 oC and P = 1 bar. (Data taken from refs 28 and 45). Ternary Systems with an Insoluble Solid Species.
Figure 11 shows an example of a ternary system in which a solid is introduced that can
contain butanol and water. The solid phase is silicalite, which is virtually insoluble in both
the aqueous phase (L1) and the organic phase (L2). Again, adding this ternary species
allows the presence of a new phase with a relatively high butanol:water ratio.
Silicalite is one of many solid species that can be used for butanol adsorption. Some have
been listed in Table 4, with their adsorption capacity for butanol. Hydrophobic materials
prevail to minimize water adsorption. The butanol adsorption capacity ranges from 4% to
22% (m/m). The loading with adsorbate is a function of aqueous butanol concentration and
temperature; however, these data are usually not available. An exception is silicalite-1,50 for
which the butanol adsorption isotherm is shown in Figure 12.51 The initial slope of this line
indicates an affinity from which the value of 2.16 for log msilicalite/aq
can be calculated. This
compares favourably with the log m org/aq
BuOH values that are reported for the solvents in Table
3. Desorption of butanol can be achieved with an auxiliary fluid phase or by temperature or
pressure shift, which requires additional isotherms, which are usually unavailable.
Adsorption data of butanol are mostly available in conjuncture with the recovery of ethanol
and acetone.31,51,54,55 However, the selectivity of the adsorption processes also is dependent
on the binding capacity of the adsorbent for water. For zeolites, the water adsorption
capacity is listed in Table 5. A decrease in the Si:Al ratio leads to a decrease in
hydrophobicity and, consequently, an increase of water adsorption up to the micropore
volume. Most m ads/aq
values in Table 5 are less favorable than the m org/aq
values in Table 3.
H 2O
However, the organic solvents will dissolve in the aqueous phase, whereas the adsorbents
will not. Regeneration of the auxiliary phase is crucial for a successful operation. For
regeneration of an adsorbent, temperature-swing or pressure-swing operations can be
applied, as well as inert-purge or displacement-purge cycles operations, where an auxiliary
species is introduced.56
Figure 8. Salt effect on water solubility in organic phase as function of salt content at T =
25 oC and P = 1 bar. (Data taken from refs 28 and 45.)
Figure 9. Schematic ternary phase diagram of butanol and water with a gaseous species
forming a third phase (example: N2).
Table 3. Butanol and Water Partition and Solubility Data in Various Organic Solvents at 25
C and 1 atm
log( m org/aq
BuOH )
sesame oil
olive oil
ethyl oleate
dibutyl phtalate
methyl laurate
dibutyl maleate
castor oil
oleyl alcohol
hexyl acetate
diethyl ether
In dodecanoic acid at 20 oC. b At 20 oC.
log( m org/aq
H2O )
s org
4 × 10-6
3.7 × 10-6
7.3 × 10-4
7 × 10-5
4 × 10-3
Figure 10. Schematic ternary phase diagram of butanol and water with a water-immiscible
liquid species (example: octanol).
2.3. Selectivity of Recovery
Fermentation of butanol can be performed in batch, fed-batch, or continuous (continuous
stirred tank reactor (CSTR) or plugflow reactor (PFR)) mode. Without going in the detail of
the fermentation processes, one can note that all systems benefit from product removal. A
general process scheme that describes a butanol production system with product recovery is
defined in Figure 13. The defined streams are the feed stream Φf, the auxiliary phase stream
Φa, the product stream Φp, and the recycle stream Φr. Depending on which operation is
described in the literature, the latter two streams can also be called permeate or raffinate,
respectively. In the literature, selectivity (Sf) is most often defined as the ratio of the butanol
and water concentration ratio in the product and feed, as shown in eq 1.
The ratios can be expressed using either mass-based concentrations (C) or in mole fractions
Sf =
[CBuOH /CH2O ]p
[CBuOH /CH2O ]f )
[x BuOH /x H2O ]p
[x BuOH /x H2O ]f
To be able to determine values for the reported recovery operations, we fix the feed at 20
g/L butanol at 25 oC, unless specifically mentioned otherwise.
2.3.1. Distillation. Distillation is the traditional method to recovery butanol from aqueous
fermentation broth. Because water is the light key component, most of the energy
consumption during distillation originates from the evaporation of the water in the feed. A
binary azeotrope is obtained at 92.7 oC. Conversion of a feed of 20 g/L butanol into an
azeotropic mixture at 1 atm leads to a selectivity of Sf = 72. It is possible to break the
azeotrope by introducing a ternary compound or by changing the pressure. The specific
energy requirement can be calculated and is a function of butanol feed concentration.59 The
performance of the distillation is directly related to the energy integration applied, because
the energy consumption determines the largest portion of the operational costs. Energy
integration options are dependent on the scale and the processing plant. In conclusion, pure
butanol can be obtained at the cost of energy and investment in equipment.
2.3.2. Liquid-Liquid Demixing. Butanol and water can form a biphasic liquid; however, at
the upper limit in fermentative concentration (~20 g/L), all butanol is still soluble in the
aqueous phase. The addition of salts might be pursued to cause a phase split. The salt
contents required in the aqueous phase to reduce the solubility of butanol to 20 g/L for
sodium chloride (NaCl), lithium chloride (LiCl), sodium bromide (NaBr), and kalium
iodide (KI) are 160, 188, 270, and 430 g/L, respectively. Liquid demixing will also provide
a butanol phase. The composition of the butanol phase, in comparison to the liquid feed
composition, leads to a selectivity of 310-750 for the salt contents that have been
mentioned. Subsequently, the salts would enter the fermenter according to Figure 13, and
then these salt concentrations would be too high for normal clostridial fermentations. A
microbial cell strives to maintain a constant intracellular environment. For larger salt
gradients over the cell membrane, more energy is spent on maintaining the cells’ internal
conditions, usually at the expense of growth or product formation. The key factor still
remains to be the overall product yield on the substrate. Nevertheless, some microorganisms live in medium- to high-saline environments, such as soda lakes, but these
organisms have not been used for butanol production.
For most other polar compounds, the effect on water and butanol demixing has not been
reported; however, it is expected that liquid-liquid demixing cannot be readily applied,
because it will require the addition of large amounts of auxiliary chemicals, which might
easily disturb the fermentation in the case of the recycle described in Figure 13. Whenever
the fermentation is not fully inhibited by the addition of large amounts of additional
chemicals, liquid-liquid demixing might be conceptually feasible.
Figure 11. Schematic ternary phase diagram of butanol and water with an insoluble porous
solid species (example: silicalite). L1 and L2 phases are located on the horizontal axis.
Table 4. Adsorbent Capacity for Butanol Recovery from Model Solutions at T = 20 oC51-53
and T = 37 oC,30 1 atm
Norit ROW 0.8
Norit W52
silica gel
Bonopore, nitrated
butanol capacity [g/g]
Figure 12. Equilibrium isotherm on silicalite from aqueous solution at T = 25 oC.51
Table 5. Water Binding Capacity in Zeolites and Silicalite at T = 25 oC57 and T = 50 oC58
log( m ads/aq
H2O )
Silicalite58 ∞
capacity [g/g]
volume [cm3/g]
Na, K-erionite
A zeolite57
NaA zeolite57
NaY zeolite57
NaX zeolite57
Figure 13. Stream definitions recovery operation.
2.3.3. Freeze Crystallization. The enthalpies required for liquid-to-solid phase transition
of either butanol or water are substantially lower than those for their respective liquid-tovapor phase transition (see Table 2). From this perspective, freeze crystallization may be
energetically more favorable than distillation.
However, process plant investment for solids handling will be significantly higher than that
for vapors, usually by a factor of 2. The energy advantage must balance the additional costs
introduced by the handling of the solid bulk water phase. No freeze recovery systems for
butanol have been described in the literature. According to Figure 3, bringing a system of
20 g/L butanol to a temperature of -20 oC will result in a butanol product stream with a
selectivity of 150.
2.3.4. Pervaporation. Pervaporation as recovery technique is a combination of membrane
permeation and evaporation.60 Pervaporation is commercially applied mainly for
dehydration of organic solvents.61 A low vapor pressure or vacuum can be used instead of
sweep gas to increase transport flux and selectivity. Membrane modules can vary from
simple sheet membranes to more-complex systems (such as tubular systems). The product
flux through the membrane and the selectivity of recovery are a function of the composition
of the aqueous phase and gas phase, membrane properties, membrane area, temperature,
and pressure. The flux through a membrane is inversely proportional to the membrane
thickness. As the membrane provides selectivity, more-selective membranes experience
relative lower product flux. Hydrophobicity of the membrane has a strong effect on
selectivity, because it limits the water flux. Silicalite, zeolite, liquid, and organic polymer
membranes are all possible considerations. In the literature, various compilations of
membrane-aided separation of butanol are available.19,62
Usually, these report batch operations with a (internal) recycle to the feed vessel and
measurement of the concentration occurs using the effluent. Therefore, the reported
selectivities should formally be denoted as Sr, with a superscript “r” instead of superscript
“f” in eq 1. In practice, most experiments are operated in fully mixed batch systems,
without an actual distinction between feed and recycle composition. A selection of
pervaporation data is shown in Table 6. Overall, the aqueous butanol concentration was in
the range of 0.37-78 /L. The selectivity is in the range of 2.7-209.
Flux is dependent on membrane thickness. The membrane thickness varied in the studies;
however, it is usually 0.025-2 mm. The total flux through the membrane varied between 3 g
m-2 h-1 and 2100 g m-2 h-1.
Flux increases with butanol feed concentration. It is reasonable to assume a standard
butanol product flux of 20-100 g m-2 h-1 to be feasible today when handling fermentation
Most membranes can be considered to be close to their performance limit, although some
membrane types (such as silicalite/silicone membranes) are considered to be improvable.63
Flux limitations can be overcome by influencing the process kinetics. The kinetics can be
influenced by process temperature or by product phase concentration. For example, the
butanol content of the product phase can be reduced using a strip gas that dilutes the system
but increases the driving force for the separation. Alternatively, a vacuum can be applied,
which leads to an increase in volume.
2.3.5. Reverse Osmosis. Traditional reverse-osmosis membranes materials dissolve in
acetone. Therefore, the presence of acetone is considered to be a main factor for the
absence of studies on osmosis as a recovery technique.62 Alternative membranes have not
been reported; however, the application of reverse osmosis as a separating technique still
remains very interesting.
2.3.6. Gas Stripping. Several gas-stripping recovery systems for butanol have been
described.1,20-23 Primarily, nitrogen is used as the stripping gas. Butanol fermentations
require anaerobic conditions, and exposure to oxygen should be avoided. Stripping can
occur in the reactor or in an external unit, the latter of which allows heating of the liquid
without disturbing the fermentation. Not only batch and fed-batch fermentation have been
conducted: some
continuous cultures also have been run on a laboratory scale, using gas stripping as a
recovery technique. In all cases, the product recovery improved the productivity and
product yield of the fermentation. An important advantage of stripping as a recovery
technique is the low risk of fouling or clogging of the auxiliary phase.1 The mass-transfer
area is determined by the gas/liquid interface, which is a function of the gas bubble size. A
compilation of operation parameters and selectivities obtained in stripping systems is
shown in Table 7.
Table 6. Pervaporation Systems and Butanol Selectivity
polytetrafluoroethylene 62
16, 62, 64
62, 65
T [oC]
2-16. 7
total flux
[g m-2 h-1]
17, 18, 62, 37-78
62, 65, 69 40-62
zeolite membrane Ge- 70
polyether block amide
62, 65
zeolite-filled PDMS
poly[1-(trimethylsilyl)- 69
thin-film silicone
silicalite-filled PDMS
18, 71
silicalite- 63
oleyl alcohol liquid 59
2.3.7. Extraction and Supercritical Extraction. The feasibility of liquid extraction of
butanol, among others, is dependent on the aforementioned partition coefficient of butanol
and water and on the solubility of the organic solvent in the aqueous phase.24,25 Usually, the
objective is to concentrate butanol in a higher-boiling solvent, enabling distillation of
butanol more efficiently than via direct distillation from the dilute aqueous solution.
Selectivity of the extraction is a function of butanol and water solubility of the solvent and
ranges from 1.2 to 4100.24,25 High selectivities are attained when extremely nonpolar
extractants are used. Although water does not readily
dissolve in such strongly nonpolar solvents, an increase in hydrophobicity adversely affects
the solubility of butanol in the solvent. This means that selective solvents have a low
capacity for butanol. For extraction with a selective solvent such as oleyl alcohol, the
selectivity at equilibrium is 105, with a capacity of only 1.8% (m/m).
The extractant will usually saturate the aqueous phase and can become toxic to the
microorganisms when the aqueous phase is recycled to the fermenter. The toxicity of an
extractant to a micro-organism can be determined experimentally, but it can be correlated to
the octanol-water partitioning coefficient.72,73
In this case, it is fortunate that very nonpolar solvents are the least toxic ones. Supercritical
CO2 (SCCO2) extraction differs from extraction only in the definition of the phase. By
reducing the pressure, the CO2 can be easily removed for recovery of the extracted
2 /aq
= 2.2, which is a good indication of the
products. Under extraction conditions, mSSCO
selectivity that one equilibrium stage can bring. Because of the multistage operation, the
selectivity is reported as Sr at a pressure of 100 bar and a temperature of 40 oC, as a
function of the butanol raffinate concentration, and ranges in value from 139 to 6020.27
Recycle concentrations were in the range of 0.7-0.14 g/L, whereas the feed contained 50
g/L. The selectivity increases with an increase in the amount of solvent stream, relative to
the feed stream. At a CO2:feed mass ratio of 1.2, the selectivity of butanol recovery is ~
400. The CO2 capacity for butanol is modest (1.6%-6.9% (m/m)).
2.3.8. Perstraction. In the literature, most applications of extraction are perstraction
operations.25,29 The membranes keep the organic and aqueous phases physically separated,
avoiding a settling compartment for the two liquid phases, avoiding contamination of the
organic phase by the cells, and reducing the toxicity of the organic phase to the cells.
Membranes do introduce an additional mass-transfer limitation. In the long run, any recycle
system, as shown in Figure 13, will operate under saturated conditions.
2.3.9. Adsorption. The selectivity of an adsorption process is dependent on the relative
binding of the adsorbent of butanol versus water. Because of the high water concentrations
and low butanol concentrations, hydrophobic material is desired. The highest adsorption
capacity for butanol in Table 4 is 22% (m/m) for activated coal. Activated coal is generally
used to remove organic contaminants from water. The regeneration of activated coal is
more cumbersome and the activated coal’s stability and homogeneity are less favorable, in
comparison to silica-based adsorbents. Because of the high affinity for butanol, even at feed
concentrations of < 1 g/L, the selectivity can be very high. The overall selectivity of
hydrophobic zeolite adsorbents is estimated to be 130-630. When working with a feed
composition of 20 g/L, the selectivity can be ~270. Desorption is an issue that has been
studied much less. It has been noted that gradual heating might allow the stepwise
desorption from silicalite-1 and therefore enable the recovery of an enriched butanol
2.3.10. Overview of Butanol Recovery Operations. Table 8 gives an overview of the
estimated selectivities of butanol recovery. Although distillation and extraction are mature
and widely applied in the chemical industry, Table 8 also contains recovery options, such as
adsorption and membrane separation, which are much less accepted, in terms of scaleup.
Still, the latter are applied on a scale larger than that required for butanol recovery. For
example, a total adsorption bed volume of 640 m3 is used for the desugarization of
molasses in an Applexion plant,75 and more than 300 000 m2 of membrane area will be
installed in the Sulaibiya wastewater treatment plant for Kuwait City.76 With a change from
hydrocarbon to carbohydrate feedstock, the chemical industry may also have to adopt such
less-mature recovery operations at very large scale.
Distillation, gas stripping, freeze crystallization, and pervaporation show a relatively low
selectivity. Liquid demixing and adsorption allow a more-selective capturing operation for
butanol. The reported selectivity is predominantly based on a single-stage equilibrium
operation. Operations such as distillation and extraction can easily be performed in
multistage contactors. Multistage equilibrium operations will allow higher selectivities,
compared to the single equilibrium stage.
Table 8 is limited to single-equilibrium-stage operations. On a mass basis, the highest
capacities are achieved for adsorption and extraction. However, for extraction, capacity is
inversely coupled to the overall selectivity of recovery. For liquid demixing, the known
salting-out agents are not compatible with the fermentation.
Based on these data, several recovery operations will be evaluated further.
2.4. Energy Requirement Estimates per Recovery Operation
The energy requirement of the recovery system contributes significantly to the operational
costs of a recovery system, as mentioned in the Introduction. The internal combustion
energy of butanol (36.2 MJ/kg) is clearly beyond the upper limit amount of energy to be
spent on recovering butanol. A target value of 10% of the combustion energy seems
reasonable, which would give an operational energy requirement limit of 4 MJ/kg. The
product-capturing step concentrates the butanol to a much smaller stream, similar to that
indicated by Figure 1. This process step is more intensive, with regard to the investment of
equipment and the energy requirement, than the subsequent purification step of the enriched
organic phase. When concentrating butanol to 50%, the capturing step consumes ~ 90% of
the energy requirement of the product recovery. For systems such as gas stripping that do
not achieve such a 50% concentration, still 75% of the energy is consumed by the capturing
step. To avoid the need for detailed product specifications, depending on too many
variables (such as byproduct and production site-specificintegration options), only the
product-capturing step is investigated in detail. It is the crucial process step when
determining the process performance. Most process energy estimates are based on process
flowsheet calculations. A simple quantitative approach to evaluate a set of recovery options
is presented here. We propose to neglect the process-specific details altogether and describe
the product-capture operation as a general steady-state flow process.
Steady-state flow processes can be described as shown in eq 2.
In the absence of a change in kinetic or potential energy, the enthalpy difference between
feed and product streams of the steady-state process is dependent on the heat and work
applied to the capture operation.77
ΔH = Q - WS
Because process-dependent characteristics are not taken into account, the heat and work are
not specified.
The course of ΔH is variable throughout the product capture operation and is dependent on
the local temperature and pressure. The recovery energy requirement is estimated using the
enthalpy extremes as illustrated in Figure 14 and eq 3. The product flow is taken to be a
specific fraction of the feed flow. The product flow is evaluated as only the mass in the feed
flow that will eventually comprise the final product flow. For simplicity, the difference in
enthalpy between the various streams are taken as a function of temperature and pressure
and not of composition. In other words, the mixing effect is neglected. Only the enthalpy
change of the product part of the feed flow will be taken into account; this is called ΔHp.
We assume the remainder of the feed flow to leave the system as recycle to the fermenter at
the original feed enthalpy level.
H p = H p (T
p,|max |
, P p,|max | ) - H p (T f , P f )
Note that energy input has been calculated from the enthalpy value of the final product
stream; it is not the actual reversible work performed on the system. Under relevant
conditions, the enthalpy is more sensitive to temperature than to pressure. In the case of
work-intensive operations that involve pressure, the energy requirement calculation is an
underestimation, because the work is assumed to be ideally applied. Butanol and water
enthalpies were calculated in Aspen Engineering Plus 12.1, using the NRTL property
method. The energy requirement for any recovery operation j is expressed per mass amount
of butanol product:
E pBuOH ( j) =
H p
x pBuOH
 M w,BuOH
Equation 1 can be used to express the product butanol composition as a function of
selectivity and feed composition.
Substitution of eq 1 in eq 4 will yield eq 5. Equation 5 expresses the energy estimated for
recovery operation j for the recovery of 1 kg of butanol from an aqueous feed mixture and
is only a function of the selectivity. Thus, the energy requirement per kilogram of butanol
in the product phase can easily be calculated using the tabled selectivities:
( j) 
H p
1 xH O
 (  f 2  1)
M w, BuOH S xBuOH
If an auxiliary phase is present, its energy change also must be taken into account. A
simpler method was applied with the assumption that the ternary species does not
experience a change in phase or pressure. Thus, the enthalpy requirement for the auxiliary
phase (ΔHpa) is estimated using its heat capacity, cap.
For convenience, the heat capacities are expressed on a mass basis, because the capacities
are expressed on a mass basis.
H a ( j) = cap (T p - T f )
The enthalpy effect of the temperature change of the auxiliary phase, combined with the
amount of butanol in the auxiliary phase (LaBuOH), given in Table 8, gives an estimate for
the energy requirement of the auxiliary phase per kilogram of butanol.
E a ( j) =
H a (j)
Table 7. Selectivity of Integrated Recovery System by Gas Stripping
mode ref
CBuOH f [g/L]
< 10
Table 8. Selectivity and Capacity Estimates Per Recovery Operationa
selectivity estimate
13, 14
liquid demixing
freeze crystallization
26, 27
gas stripping
1, 20–23
24, 25
1.2 - 4100
No auxiliary phase involved. b Low capacity at high selectivity.
capacity [kg/kg]
Table 9. Energy Requirement Estimates per Kilogram of Butanol
Energy Requirement
Contribution [MJ/kg]
operation, j
gas strippinga
Gas phase product.
No heat is assumed to be lost to fermentation broth.
vacuum operation is assumed at 100% efficiency.
Work of
Not for the auxiliary phase but for
heating of the liquid feed to the recovery temperature.
Using oleyl alcohol.
The heat
capacity of the adsorbent is assumed to be 1 kJ kg-1 oC-1.
Figure 14. Enthalpy during recovery.
Table 10. Energy Requirement of Butanol Recovery Systemsa
Energy Requirement (MJ/kg)
(from ref 59)
(from ref 31)
stripping 24
(from ref 14)
gas stripping
extraction/perstraction 9
Results have been normalized per kilogram of recovered butanol.
Acetone, 1-butanol,
and ethanol.
Both contributions are reported separately in Table 9, to provide a clear picture of the
energy requirements of the recovery options. Low-energy-requirement estimates can be
observed for selective adsorption processes and pervaporation processes without preheating
the feed stream. It is necessary to have a highly selective system and to avoid high energy
demands in the auxiliary phase to remain below the already-mentioned target value of 4
MJ/kg butanol. Table 9 lists the energy requirements for each recovery operation and
conditions. The contribution of auxiliary phase to the energy consumption can be
considerable and corresponds to a large investment in heat-transfer equipment. The
operations listed with footnote (c) correspond to processes mentioned in the literature,14 as
shown in Table 10.
The applied method ranks the recovery operations in similar order to the predictions in the
literature on process estimates (shown in Table 10). The short-cut method results shown in
Table 9 underestimate the required energy, when compared to Table 10. This is expected as
the short-cut method evaluates the internal enthalpy change and not the actual heat and
work performed on the system. The processes with energy requirements below the target
value of 4 MJ/kg are, according to this short-cut method, pervaporation and adsorption,
provided that high selectivities are achieved.
2.5. Conclusions
The thermodynamic properties of binary and ternary butanol-water mixtures lead to a wide
range of options available for the recovery of butanol from dilute aqueous solution, by
varying temperature, pressure, or composition. Recovery options that have not yet been
explored include freeze crystallization and liquid-liquid demixing. From an energy
requirement perspective, both operations are interesting.
Complete process designs for the recovery of butanol are scarce. Most operational data
have been reported only for model systems using batch recovery. The available recovery
data have been ranked by the selectivity of the recovery of butanol versus water. Highly
selective processes are adsorption-based recovery processes using nonpolar adsorbents and
some extraction systems. Selective extraction systems show a very low capacity for
butanol. Ranking of the energy consumption of the recovery alternatives has been achieved
using a short-cut model, which only takes into account the enthalpy change during the
recovery operation. According to this assessment, the most attractive recovery options are
adsorption-and pervaporation-based recovery operations.
2.6. Acknowledgment
This project is financially supported by The Netherlands Ministry of Economic Affairs and
the B-Basic partner organizations ( through B-Basic, a public-private
NWOACTS programme (ACTS ) Advanced Chemical Technologies for Sustainability).
2.7. Nomenclature
C = mass concentration [g/L]
cp = heat capacity [J g-1 K-1]
E = mass-specific energy requirement [MJ/kg]
F = molar flow rate [mol/s]
G = Gibbs energy [J/mol]
H = specific enthalpy [J/mol]
L = capacity [kg/kg]
Mw = molar mass [g/mol]
m = distribution coefficient
P = pressure [atm]
Q = mass-specific heat [J/kg]
s = solubility [g/L]
S = selectivity
T =temperature [oC]
Ws = mass-specific shaft work [J/kg]
x = mole fraction in liquid phase
y = mole fraction in vapor phase
Φ = stream volume flow rate [m3/h]
b = boiling
BuOH = butanol
comb = combustion
fus = fusion
H2O = water
j = process operation
m = melting
vap = vaporization
∞= infinite
a = auxiliary phase
aq = aqueous phase
f = feed
org = organic phase
p = permeate/product
r = recycle/retentate
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Chapter 3: Exploring the potential of recovering 1-butanol from aqueous solutions by
liquid demixing upon addition of carbohydrates or salts
BACKGROUND: Fermentative production of 1-butanol yields dilute aqueous solutions.
Recovery of the butanol from these solutions is most commonly performed by energyintensive distillation. This work investigated the liquid-liquid (L-L) phase behavior of
mixtures of butanol and water to explore the potential of using L-L phase separation as a
recovery possibility for 1-butanol. The phase behavior is preferably influenced by
compounds already present in the fermentation, such as carbohydrates and salts.
RESULTS: The L-L phase equilibria of butanol and water were determined in the presence
of glucose, fructose, sucrose, NaCl, LiCl and CaCl2. The aqueous and organic phase split is
more pronounced in the presence of salts than in the presence of carbohydrates. Demixing
is achieved with about 0.3 kg salt kg−1 aqueous phase containing 40 g of butanol.
CONCLUSION: Operation of L-L based recovery using salts or carbohydrates requires
extreme concentrations of those compounds. For feed material containing 40 g kg−1
butanol, the tested carbohydrates do not influence the phase equilibria sufficiently to allow
butanol separation. Fermentative butanol concentrations up to 70 g kg−1 are required to
create an effective L-L phase split. The remaining residual aqueous carbohydrate solution
might be used as feed for a following fermentation.
Published as: Exploring the potential of recovering 1-butanol from aqueous solutions by
liquid demixing upon addition of carbohydrates or salts, Arjan Oudshoorn, Marjolein
C.F.M. Peters, Luuk A.M. van der Wielen and Adrie J.J Straathof, J Chem Technol
Biotechnol 2011; 86: 714-718
3.1. Introduction
The dominant transportation fuels today are gasoline, diesel and kerosene. Sustainability
concerns are a major driving force behind the development of technologies enabling
renewable feedstock conversion into energy carriers,1–3 in particular into ethanol.4
However, higher alcohols, e.g. 1-butanol, provide better product characteristics, such as
higher energy content on a mass basis. 1-butanol can be produced by Clostridia using
acetone–butanol–ethanol (ABE) fermentation.5 Due to product inhibition, the overall
concentration of the solvents in the production broth reaches only around 20 g kg−1.6–8 This
total solvent concentration might be improved significantly by organism choice, because
the highest butanol tolerance has been reported to be 48 g kg−1.9 Owing to the product
concentration, separation and purification of the dilute aqueous stream by traditional
recovery techniques, such as distillation, require a relatively high energy input when
compared with the energy content of butanol.10 Alternative recovery techniques are
pervaporation, adsorption, and extraction.11 All these techniques use the introduction of a
new product-enriched phase in order to separate the product from the bulk fermentation
broth. A mostly unexplored technique, which is more direct, is liquid demixing. Butanol
and water are not fully miscible under most conditions. At ambient conditions, 74 g kg−1
butanol dissolves in water. The liquid–liquid equilibria can be influenced by pressure,
temperature and chemical composition.12–14 We will not explore temperature and pressure
influences on the phase equilibria here, but will focus only on chemical composition. This
work explored the use of polar solutes, some of which were already present in the
fermentation process, to influence the chemical composition of a two-phase organic and
aqueous system containing butanol. The feed material used for fermentations contains
mostly carbohydrates, but also some salts are added as nutrient. Salts are further introduced
in the form of acid and base dosing for pH control. The aim is to show the possibilities and
(operating) boundaries for an integrated recovery process using L-L phase behavior as the
separating technique. To that end liquid–liquid phase equilibria between butanol, water and
a ternary species, carbohydrates as well as various salts are determined by batch
equilibrium experiments. The equilibria are used to calculate the potential of using the
various chemical species to perform a L-L phase separation process, as is depicted in Fig. 1.
Figure 1. General process scheme for 1-butanol production using L-L separation based
product recovery.
3.2. Experimental section
D-glucose (>99.0%), D-fructose (>99.0%), sucrose, CaCl2•2 H2O (>99.5%) and,
phosphoric acid (85%) were supplied by Merck (Schiphol-Rijk, the Netherlands); NaCl
(99.7%) and LiCl (>99.6%) by Baker (Deventer, the Netherlands); and 1-butanol (>99.5%)
by Acros (Geel, Belgium).
3.2.2.Miscibility experiments to determine phase equilibria
The equilibrium experiments were performed by adding a set mass of water, butanol and a
carbohydrate or salt species into a 250 mL vessel. The mixture was stirred magnetically for
about 4 h and then transferred into a separating funnel. The funnel was placed inside an
oven kept at 25 ◦C. After approximately 24 h of settling, the two liquid phases were
separated. Their mass was determined and they were stored at 5 ◦C for further analysis.
3.2.3.Analytical methods
A high performance liquid chromatograph (HPLC,Waters,Milford, MA, USA) was used to
determine the butanol concentration in all samples, with a 2414 Refractive Index (RI)
detector (Waters, Milford, MA, USA), a 515 HPLC pump and a 717plus autosampler. A
BioRad Aminex (Biorad, Hercules, CA, USA) HPX- 87H (7.8 mm, 300 mm) column was
used, with a dilute phosphoric acid solution as mobile phase, to determine the aqueous and
organic phase concentrations of fructose and glucose. Sucrose concentrations were
determined by an enzymatic assay, using a sucrose/D-glucose kit from R-Biopharm
(Darmstadt, Germany).
Absorption was measured at 340 nm on a Tecan (MTX Lab Systems, Vienna, VA, USA)
GENios, with a costar3631, flat-bottom, non-treated microtiter plate. ICP-OES was used to
determine the salt concentrations in the organic phase, using a PerkinElmer (Waltham, MA,
USA) Optima 5300dv; all samples were diluted 100 times and were brought to acidic
conditions with 0.5 mol L−1 nitric acid. Non-volatile residues were measured
gravimetrically, after overnight drying in an oven, to determine the salt content in aqueous
samples. Density measurements were carried out on a vibrating-tube Anton Paar GmbH
(Ostfildern, Germany), DMA48, to enable the use of mass balances (over the species and
phases) to calculate the water content of each phase.
3.3. Results
All results are at 25 ◦C and 1 atm. The mass fractions for the two phase systems of butanol,
water and dissolved carbohydrate are shown in Tables 1 to 3 for glucose, fructose and
sucrose, respectively. Comparable trends are seen for the three carbohydrates.
Table 1. Mass fractions in liquid-liquid phase equilibria of butanol, water and glucose
Aqueous phase
H2 O
Organic phase
Table 2. Mass fractions in liquid-liquid phase equilibria of butanol, water and fructose
Aqueous phase
Organic phase
Table 3. Mass fractions in liquid–liquid phase equilibria of butanol, water and sucrose
Aqueous phase
H2 O
Organic phase
Figure 2. Ternary phase diagram for water, butanol and glucose two phase system.
Compositions are given for mass fractions. Tie-lines are shown between the organic and
aqueous phase equilibria points.
Table 4. Mass fractions in liquid–liquid phase equilibria of butanol, water and NaCl
Aqueous phase
Organic phase
Table 5. Mass fractions in liquid–liquid phase equilibria of butanol, water and LiCl
Aqueous phase
Organic phase
In the concentration range studied, addition of the carbohydrates decreases the
butanol/water ratio in the aqueous phase and the water/butanol ratio in the organic phase by
values in the range 30–70%. The equilibrium data for the glucose–butanol–water system
are shown graphically in a ternary phase diagram (Fig. 2). The L-L mass fractions for
butanol, water and salt solutions (NaCl, LiCl and CaCl2, respectively), are presented in
Tables 4 to 6. The salt mass of CaCl2•2H2Ohas been corrected for water of hydration. The
water of hydration was added to the total amount of water present in the system for mass
balance calculations when determining water content. On mass basis, the decrease of
butanol solubility in water is more pronounced for the addition of salts than for the addition
of the carbohydrates investigated, according to Fig. 3. The addition of CaCl2 or LiCl leads
to a local maximum of the organic phase butanol mass fraction of 0.89 and 0.82,
respectively, at ~ 0.002 mass fraction of ternary species, according to Fig. 4. For ~ 0.001
mass fraction NaCl in the organic phase there is a local maximum that is less clear, but also
at 0.89 mass fraction butanol. The addition of carbohydrates leads to an increase of butanol
mass fraction in the organic phase from ~ 0.78 to ~ 0.90, (see Fig. 5).
Table 6. Mass fractions in liquid–liquid phase equilibria of butanol, water and CaCl2
Aqueous phase
Organic phase
Figure 3. Butanol fraction versus salt and carbohydrate mass fraction in the aqueous phase.
3.4. Discussion
L-L based phase separation, when using a dilute aqueous butanol phase, requires accurate
determination of liquid phase compositions at low concentrations, because experimental
and measurement errors can significantly influence the recovery performance calculations.
Butanol–water–NaCl compositions of the aqueous phase from this study agree with those in
the literature.15,16 The organic phase compositions show some stronger deviations from
those in the literature. This can be explained because the organic phase compositions are
more difficult to determine accurately than the aqueous compositions, in particular with
respect to butanol content. Butanol and salt content of the organic phase have been
calculated using species mass balances by some authors,16 to avoid direct butanol
measurements of the organic phase.
Figure 3 shows that in saturated aqueous solutions containing a desired low mass fraction
of butanol, about three times more carbohydrate than salt must be present. When
minimizing the amount of ternary component necessary to perform phase separation, NaCl,
LiCl and CaCl2 are used preferably to glucose, fructose and sucrose.
Figure 4. Butanol fraction and salt mass fraction in the organic phase (Δ, CaCl2) (◊, LiCl)
( , NaCl).
Figure 5. Butanol fraction and carbohydrate mass fraction in the organic
phase (+, glucose) (×, fructose) (◦, sucrose).
Figure 6. Butanol recovery of the feed, versus the amount of auxiliary species added to the
feed (◊, LiCl) ( , NaCl) (Δ, CaCl2).
Figure 7. Butanol recovered per amount of auxiliary material added, versus the amount of
auxiliary species added to feed (◊, LiCl) ( , NaCl) (Δ, CaCl2).
For salt based recovery, the selectivity of separation, as defined in Chapter 2 equation 1,
when starting with a feed containing 20 g/L butanol, the maximum selectivity are 388, 208
and 412, for NaCl, LiCl and CaCl2, respectively. However, at those levels most of the
butanol will be in the aqueous phase and only a small, highly enriched, organic phase will
be present.
Figure 6 shows the amount of ternary species necessary when performing a demixing
operation on an aqueous 40.8 g kg−1 butanol solution. The results are obtained by
calculating the masses of the species in organic and aqueous phase at equilibrium.
Recovery of 70 to 80% butanol requires addition to the feed of about 400 g kg−1 of NaCl,
LiCl or CaCl2. The costs of the cheapest ternary compound (NaCl) may be ~0.15 $ kg−1, so
if 10 kg NaCl is spent to produce 1 kg of butanol, which may be worth~1.8 $ kg−1, this
approach is unaffordable because there will be other major costs. Besides, using such large
proportions of salts in industry is inconvenient.
The capacity of the auxiliary material, which is defined here as the amount of butanol
recovered per amount of auxiliary material introduced, is shown in Fig. 7 as a function of
the amount of auxiliary species that has to be added to reach the new equilibrium. Similar
mass ratios are found for many separation systems using an auxiliary phase.10
Unfortunately, the effect is less pronounced for the carbohydrates than for the salts. An
important incentive for using carbohydrates in a L-L based recovery system is that the
carbohydrates can be consumed upstream. This can be accomplished by recycling the
remaining aqueous phase as feed for the fermentation section, see also Fig. 1. Continuous
feeding might be applied to prevent that the microorganisms operate at too high sugar
concentrations, and storage tanks may have to be implemented. To prevent microbial
contamination, the location of the sterilization of the sugar solutions needs careful
As long as the carbohydrate consumption matches the amount required for liquid phase
split, no additional input to the process is necessary. However, on a mass basis
carbohydrate consumption is only about twice the butanol production, as butanol
production yield on glucose has a maximum of 0.42 g g−1.17 A typical consumption of
carbohydrate would thus be ~40 g kg−1 water, which is much lower than the amount of
carbohydrate required for liquid demixing. Carbohydrates that have a stronger demixing
effect would be required for this approach to be effective. A compound like sodium
gluconate, which combines carbohydrate and salt properties, might behave much better, but
is too expensive to be converted into butanol. The viscosity, density and surface tension of
a carbohydrate mixture complicates the process further.
Growth of microorganisms under saline conditions has been shown to be possible under
severe conditions. Growth up to 4mol L−1 Na+ has been reported.18 However, this has not
been combined with the production of butanol or butanol-producing microorganisms. Also,
sugar tolerant microorganisms like yeasts can grow on media containing up to 220 g kg−1
total sugar.19 A butanol-producing metabolic pathway might be introduced into such a
microorganism, but obtaining a sufficiently efficient strain would be very complicated.
As mentioned, the influence of the carbohydrates investigated on the L-L phase equilibria is
less pronounced than for the salts, so the recovery is lower. When starting with 70 g kg−1
butanol, addition of 150 g of carbohydrate will provide an organic phase containing only
5% of the butanol present.
It might be interesting to study this approach for other alcohols that can be produced by
fermentation, such as isobutanol and pentanols20,21 to see if a stronger effect on the L-L
phase behaviour can be found.
3.5. Conclusions
Demixing aqueous 1-butanol into an aqueous and butanol phase requires addition of large
amounts of salts, exceeding 250 g kg−1 total solution, in order to recover 70% of the
product, for aqueous phase containing 40 g butanol per kg. Demixing might become useful
for recovery of butanol from aqueous fermentation solution if the butanol concentrations
would be almost at saturation, i.e. ~74 g kg−1.
3.6 Acknowledgements
This project is supported financially by the Netherlands Ministry of Economic Affairs and
the B-Basic partner organizations ( through B-Basic, a public–private
NWO-ACTS program (ACTS = Advanced Chemical Technologies for Sustainability).
3.7 References
1 Antoni D, Zverlov VV and Schwarz WH, Biofuels from microbes. Appl Microbiol
Biotechnol 77:23–35 (2007).
2 Petrus L and Noordermeer MA, Biomass to biofuels, a chemical perspective. Green
Chem 8:861–867 (2006).
3 Straathof AJJ, Panke S and Schmid A, The production of fine chemicals by
biotransformations. Curr Opin Biotechnol 13:548–556 (2002).
4 Hill J, Nelson E, Tilman D, Polasky S and Tiffany D, Environmental, economic,
andenergetic costs andbenefits of biodiesel andethanol biofuels. Proc Natl Acad Sci USA
103:11206–11210 (2006).
5 Dürre P, New insights and novel developments in clostridial acetone/butanol/isopropanol
fermentation. Appl Microbiol Biotechnol 49:639–648 (1998).
6 Lee SY, Park JH, Jang SH, Nielsen LK, KimJ and Jung KS, Fermentative butanol
production by clostridia. Biotechnol Bioeng 101:209–228 (2008).
7 Papoutsakis ET, Engineering solventogenic clostridia. Curr Opin Biotechnol 19:420–429
8 Qureshi N and Blaschek HP, Economics of butanol fermentation using hyper-butanol
producing Clostridium beijerinckii BA101. Food Bioprod Process 78:139–144 (2000).
9 Ruhl J, Schmid A and Blank LM, Selected Pseudomonas putida strains able to grow in
the presence of high butanol concentrations. Appl Environ Microbiol 75:4653–4656 (2009).
10 Stark D, Jaquet A and von Stockar U, In-situ product removal (ISPR) in whole cell
biotechnology during the last 20 years. Adv Biochem Eng Biotechnol 80:149–175 (2003).
11 Oudshoorn A, van der Wielen LAM and Straathof AJJ, Assessment of options for
selective 1-butanol recovery from aqueous solution. Ind Eng Chem Res 48:7325–7336
12 Sorensen JM, Liquid-liquid equilibrium data collection; binary systems. Dechema
Chemical Data Series V:236–237 (1979).
13 van Berlo M, Ottens M, Luyben KCAM and van der Wielen LAM, Partitioning
behavior of amino acids in aqueous two-phase systems with recyclable volatile salts. J
Chromatogr B: Anal Technol Biomed Life Sci 743:317–325 (2000).
14 Al-Sahhaf TA and Kapetanovic E, Salt effects of lithium chloride, sodium bromide, or
potassium iodide on liquid-liquid equilibrium in the system water plus 1-butanol. J Chem
Eng Data 42:74–77 (1997).
15 Li ZC, Tang YP, Liu Y and Li YG, Salting effect in partially miscible systems of nbutanol water and butanone water. 1. Determination and correlation of liquid-liquid
equilibrium data. Fluid Phase Equilib 103:143–153 (1995).
16 de Santis R, Marrelli L and Muscetta PN, Liquid-liquid equilibria in water-aliphatic
alcohol systems in the presence of sodium chloride. Chem Eng J 11:207–214 (1976).
17 Tashiro Y, Shinto H, Hayashi M, Baba SI, Kobayashi G and Sonomoto K, Novel highefficient
saccharoperbutylacetonicum N1-4 (ATCC 13 564) with methyl viologen. J Biosci Bioeng
104:238–240 (2007).
18 Sorokin DY, Banciu H, van Loosdrecht MCM and Kuenen JG, Growth physiology and
competitive interaction of obligately chemolithoautotrophic, haloalkaliphilic, sulfuroxidizing bacteria from soda lakes. Extremophiles 7:195–203 (2003).
19 Morimura S, Ling ZY and Kida K, Ethanol production by repeated batch fermentation at
high temperature in a molasses medium containing a high concentration of total sugar by a
thermotolerant flocculating yeast with improved salt-tolerance. J Ferment Bioeng 83:271–
274 (1997).
20 Cann AF and Liao JC, Pentanol isomer synthesis in engineered microorganisms. Appl
Microbiol Biotechnol 85:893–899 (2010).
21 Atsumi S, Hanai T and Liao JC, Non-fermentative pathways for synthesis of branchedchain higher alcohols as biofuels. Nature 451:6–U13 (2008).
Chapter 4: Adsorption equilibria of bio-based butanol solutions using zeolite
1-Butanol can be produced by clostridial fermentations with acetone and ethanol as byproducts. The butanol can be present up to ~20gL-1 depending on process conditions and
microbial strain. The high-silica zeolite CBV28014 has been proven to adsorb butanol
selective over water, while showing higher affinity for butanol than for acetone and ethanol.
Multi-component acetone–butanol–ethanol (ABE) adsorption on CBV28014 has been
modeled using a single site extended Langmuir adsorption model and the ideal adsorbed
solution (IAS) theory model. The IAS model describes multi-component adsorption of
ABE in synthetic mixtures and ABE in filtered fermentation broth by CBV28014 more
accurately than the single site extended Langmuir model.
Bioseparations Adsorption Bioresources Biofuels Butanol Zeolite
Published as: Adsorption equilibria of bio-based butanol solutions using zeolite, Arjan
Oudshoorn Luuk A.M. van der Wielen, Adrie J.J. Straathof. Biochemical Engineering
Journal 48 (2009) 99-103.
4.1. Introduction
Production of biofuels is currently one of the key strategies in developing a sustainable
economy. In all probability a whole range of biofuels, from hydrogen and ethanol to
biodiesel, will appear on the transportation fuel market in the first half of this century [1].
In comparison to ethanol, e.g., 1-butanol is an attractive biofuel candidate as it has positive
fuel characteristics such as carbon chain length, volatility, polarity and combustion value.
Butanol can be biologically derived from carbohydrates by means of clostridial
fermentations, with acetone and ethanol as principle by-products [2]. The highest butanol
concentration obtained in clostridial fermentations is around 20 g L-1 [3]. This limitation
occurs because butanol increases the permeability of the microbial cells’ membrane.
Maintaining homeostasis becomes increasingly more difficult at higher butanol
concentrations, leading to a complete stop of all microbiological activity. Microorganisms
with a higher butanol tolerance are being discovered and developed, but elimination of the
inhibiting characteristics of butanol is unlikely. Therefore, efficient recovery of butanol
from dilute aqueous solutions is a prerequisite for development of a biobutanol production
process. Adsorption has been shown to be a promising recovery technique [4]. Adsorption
should allow separation of the butanol from the bulk aqueous fermentation broth, so that
further downstream processing operations need to be performed on a relatively small
amount of organic product phase only; thus providing the basis for energy efficient
recovery. Hydrophobic adsorbents potentially show the desired high selectivity for butanol
over water. Zeolitic hydrophobic sorbents have additional positive features such as stability,
homogeneity and low heat capacity. The hydrophobicity of zeolites increases with
increasing SiO2/Al2O3 ratio, with silicalite containing no alumina. Silicalite-1 (ZSM-5
structure) has been shown to adsorb low amounts of water [5]. Commercially available
zeolites with high silica over alumina content have been shown to adsorb organic
components selectively over water [6]. Adsorption of butanol from fermentation broth has
been reported for various zeolites for binary mixtures and some multi-component systems,
e.g., with ethanol and acetic acid [7,8]. However, there is a lack of data on butanol
adsorption by high silica zeolites in the presence of acetone and ethanol. Also, the butanol
adsorption behavior of commercial high silica zeolites from aqueous mixtures or
fermentation broth is lacking. Therefore, we have determined the adsorption of butanol and
water by three structurally different commercially available high-silica zeolites. The
competitive adsorption of acetone, butanol and ethanol from aqueous mixtures and
fermentation broth has also been measured and mathematically modeled.
4.2. Materials and methods
4.2.1. Materials
The powdered zeolites, CBV28014, CBV901 and CBV811C-300 (abbreviated to CBV811)
were from Zeolyst International, Conshohocken, PA, USA. All zeolites were calcined at
600 oC for 8 h. Further details on the zeolites are shown in Table 1.
Table 1: Manufacturer’s specifications of zeolite adsorbents and pore volume estimates.
Zeolite type
Surface area Estimated
pore volume
Beta (BEA)
The pore volumes of the specific zeolites listed in Table 1 are not reported in the literature.
For silicalite-1 (alumina free ZSM-5, MFI framework) the pore volume is 0.19 cm3/g [5].
The MFI framework has an intersecting channel system with both channel systems being
10-ring channels (0.53 nm × 056 nm) and (0.55 nm × 0.51 nm) [9].
Beta-type zeolites consist of faulted intergrowth of two separate 3-dimensional 12-ring pore
structures. The two structures have a pore structure of channels with (0.73 nm × 0.60 nm)
and (0.56 nm × 0.56 nm) dimensions [9].
The pore volume for CBV901 has been estimated at 0.24 cm3/g [10], but this value is
surprisingly low in comparison to reported volume for NaY zeolites. The FAU-type
framework of the zeolite has a large void volume, ~50%, which is easily accessible. FAUtype zeolites have 12-ring pore openings and a 3-dimensional channel system (0.74 nm ×
0.74 nm) [9].
1-Butanol (Acros, purity 99.5%), acetone (Merck, purity 99.9%) and ethanol (Merck, purity
99.9%) were used. Samples from two clostridial fermentations were obtained from the
Wageningen University and Research Centre. One sample was directly taken from the main
fermentation broth. The other sample was filtered fermentation broth and contains no
microbial cells. The fermentation broth contained 2.33, 9.02, 0.25, 0.45 g L-1 acetone,
butanol, ethanol and butyrate, respectively, according to gas chromatography (GC)
analysis. The filtrate contained 1.72, 4.84, 0.14, 0.5 g L-1 acetone, butanol, ethanol and
butyrate. The pH of the fermentation broth and filtrate were 5.16 and 5.36, respectively.
4.2.2. Experimental methods
Experiments were carried out at temperatures in the range of 22–25 oC. Gas phase equilibration experiments
Gas phase equilibrium experiments were performed in a closed desiccator. A known
amount of the calcined zeolite was placed next to an amount of liquid butanol or water. The
mass increase of the zeolite after 72 h was used to calculate its adsorption of the species
involved. The adsorption of the sorbate via gas phase equilibrium was measured after 90 h.
It was verified equilibrium had been reached because no deviation in adsorption was seen
after 264 h. Liquid phase equilibration experiments
Liquid phase adsorption equilibrium experiments were carried out in closed stirred vessels
of 35–40 mL. The vessels contained known amounts of butanol, acetone, ethanol, water and
adsorbent. Typically 0.25–1 g adsorbent was used. After 48 h equilibration, acetone,
butanol and ethanol liquid phase concentrations were determined. The adsorption q of a
sorbate (i) on a zeolite (z) was calculated by means of mass balance as shown in Eq. (1).
Over the concentration range the density of the mixtures varies by less than 0.3%.
Therefore the density of the liquid phase was considered to be constant at the density of
pure water. The volume was calculated using the initial total mass of acetone, butanol,
ethanol and water. Competitive water adsorption by the zeolite material could not be
measured by us and was not taken into account. The error that imposes on the calculation of
the amount of adsorption of the other compounds is up to 3 times smaller than the error in
the adsorption due to the GC measurements. The accuracy of a calculated adsorption is a
function of the ratio between the aqueous concentration and available solid zeolite material
qi =
(Ci,0 - Ci )  V0
mz ,0
The measured adsorption isotherms were modeled using a Langmuir-type equation for
single site adsorption, neglecting water adsorption. For multi-component mixtures
containing (j) species the Langmuir model for adsorption of component (i) is shown in Eq.
qi =
q m,i  K i  Ci
 Cj
The ideal adsorbed solution (IAS) theory has originally been developed to describe gas
phase adsorption of volatile components. The IAS model has been modified to describe
liquid phase adsorption [11]. IAS theory makes use of the spreading pressure (π), defined as
the difference between the interfacial tension of the pure solvent–solid interface and the
solution–solid interface area (A), given for a specific amount of solid phase. The following
Gibbs relation (Eq. (3)) shows this relation [11]. The superscript o denotes that the system
is seen as a single-solute system. IAS calculations require the single solute and solvent
properties to be expressed on a molar basis
Table 2: Gas phase single-component equilibrium adsorption [g g-1] at 25 oC.
 A
R T
0.061 ± 0.009
0.27 ± 0.015
0.47 ± 0.012
0.12 ± 0.011
0.34 ± 0.004
0.37 ± 0.007
The IAS model uses the single-solute adsorption at equilibrium, qio, to calculate the
spreading pressure for each solute i. The basis of the IAS model is given by Eq. (4) and
links the (total) solvent concentration in a multi-component mixture to the actual solid
phase loading in a multi-component mixture. Cio(π) is such, that solute i adsorbs singly
from solution, at the same temperature and spreading pressure as the mixture does. The
adsorbed species in the solid phase can be calculated for any liquid phase composition. For
the complete derivation of Eq. (4) we refer to the original authors [11]
CT  x i = Cio ( )  zi
The IAS model equations and variables can be solved with the following additional
relations. The fraction of species i in the solid phase, zi, is calculated from the total
adsorption of the solid phase qT, and is shown in Eq. (5)
qT =
q io
The spreading pressure for all species should be equal at equilibrium. Furthermore, the sum
of each fraction equals one, as shown in Eq. (6)
z =  x
The solid phase loading can be calculated for any multi-component adsorption system,
using the relations given in (5) and (6), by simultaneously solving Eqs. (3) and (4) for all
components. After solving the equations, the results can be converted from molar to mass
basis. Analytical methods
Determination of the concentrations of acetone, butanol and ethanol in the aqueous phase
was performed by means of GC, Thermo Electron Corporation, model GC-Focus, with an
autosampler, AS3000. The GC column was Innowax 19091N-133 (30 m × 0.20 mm, with
coating of 0.25 μm) from Agilent Technologies Inc. Mobile carrier gas phase through the
column was helium at 6 mL/min. The temperature of the column was at 70 oC for 1 min and
increased to 130 oC with a heating rate of 10 oC min-1 after which the column was kept at
130 oC for 5 min. A flame ionization detector was used. 1-Pentanol was added to the
samples as internal standard.
4.3. Results
4.3.1. Single-component equilibration
Gas phase equilibrium experiments were carried out in triplicate to determine the
equilibrium adsorption for single components. The results are shown in Table 2. Like
expected, CBV28014 shows the lowest amount of adsorption of water [7]. The measured
adsorption is slightly higher than the literature values, which are 0.046–0.050 g g-1 for
adsorption of water on ZSM-5 zeolite material [5,7,12]. CBV811 and CBV901 show a high
adsorption capacity for both water and butanol. The adsorptions show a positive correlation
with the pore volumes of Table 1. CBV28014 seems to be far more selective than the other
two zeolites when directly comparing the single-component adsorption capacities, with a
butanol capacity of 0.11 g g-1 to a water capacity of 0.06 g g-1.
Fig. 1. Experimental adsorption of butanol on CBV28014 (♦), CBV901 (▲) and CBV811
(○) at 25 oC, with full lines for the Langmuir model.
4.3.2. Liquid phase equilibration of binary mixtures
The adsorption of butanol from aqueous solution is shown in Fig. 1. CBV28014 has the
highest affinity for butanol at low concentrations. CBV811 and CBV28014 have similar
butanol adsorption over a large concentration range, and a similar maximum adsorption
capacity, but CBV901 shows an approximate 25% higher capacity for butanol, 0.16 g g-1 at
butanol concentrations exceeding 7g L-1, maybe due to its higher internal volume. For
CBV811 and CBV901 the observed adsorptions are far below the single-component
butanol adsorptions of Table 2. This suggests that severe competition by water may occur.
In contrast, for CBV28014 the single-component butanol adsorption is already reached with
~1g L-1 butanol in water, suggesting that competition by water does not play an important
role for this case. Therefore, CBV28014 is used subsequently.
For CBV28014, comparison of adsorption isotherms of different solutes shows a decrease
in affinity for the sorbate with an increase in polarity from the sorbate, from butanol to
acetone to ethanol, see Fig. 2. This is in line with the hydrophobicity of the zeolite and
might facilitate consecutive desorption in a column operation [6,13]. The behavior of
butyric acid is not in line with this. Butyric acid is slightly more polar than butanol but
shows a higher affinity. Specific interactions between butyric acid and the zeolite might
play a role.
A two parameter Langmuir equation was fitted to the single solute isotherms data using the
non-linear Marquardt method to a two parameter Langmuir isotherm. For butyric acid the
data were transformed to a linear relationship [1/q vs. 1/C] and fitted by the least square
method. Figs. 1 and 2 show that acceptable fits were obtained. The fitted Langmuir
parameters in Table 3 indicate that for CBV28014 the affinities are widely different for the
different solutes, but the maximum amounts of adsorbed species are in a relative narrow
Fig. 2. Experimental adsorption isotherms at 25 .C for acetone (♦), butanol (▲), butyric
acid (●) and ethanol (■. on CBV28014, with full lines for the Langmuir model.
4.3.3. Liquid phase equilibration of ternary and higher mixtures
A limited number of equilibrium experiments with CBV28014 and aqueous mixtures of
acetone, butanol and ethanol was performed in relevant concentration ranges, in order to be
able to describe competitive adsorption. Table 4 shows that ethanol adsorption is very low
at these conditions, <0.003 g g-1, which is within the measuring sensitivity of the
experimental set-up. The acetone data, when compared to Fig. 2, clearly indicate that
competition by butanol occurs.
Table 3: Langmuir parameters.
K (L g-1)
3.14 ± 0.94
1.68 ± 0.18
qm (g g-1)
0.168 ± 3.6 × 10-3
0.126 ± 7.1 × 10-3
Butyric acid
42.8 ± 6.6
1.65 ± 0.42
0.34 ± 0.11
139 ± 30
0.118 ± 2.6 × 10-3
0.121 ± 5.9 × 10-3
0.093 ± 7.2 × 10-3
0.130 ± 2.5 × 10-3
Table 4: Measured equilibrium adsorption of synthetic mixtures of acetone, ethanol and
butanol by CBV28014 at 22–25 oC.
Aqueous concentration (g L-1)
Adsorbed species (g g-1)
Table 5: Measured equilibrium adsorption of acetone, ethanol, butyric acid and butanol
from filtered medium (first two entries) and fermentation broth (last two entries) by
CBV28014. Ethanol adsorption is too low for accurate determination at 22–25oC.
Adsorbed species (g g-1)
Dissolved concentration (g L-1)
Acetone Ethanol Butanol Butyric
Acetone Ethanol
Butanol Butyric
-0.0000921 0.098
Negative values due to too small aqueous concentration differences to calculate
Fig. 3. Parity plot multi-component Langmuir model predictions vs. experimental butanol
adsorption. (■) Butanol adsorption from synthetic ABE mixture by CBV28014. Butanol
(▲) adsorption from fermentation broth as shown in Table 5.
Competitive adsorption also occurred when using fermentation broth and filtered
fermentation broth (Table 5). In this case, butyric acid was also included.
Adsorption of butanol and acetone and butyric acid from fermentation broth and cell free
filtered broth behave similarly. Butyric acid adsorption at these conditions is approximately
0.005 g g-1. The results indicate that butyric acid displaces butanol. To show this more
clearly, using the parameters in Table 3, multi-component adsorption of acetone, butanol,
Fig. 4. Parity plot IAS model predictions vs. experimental butanol adsorption. (■) Butanol
adsorption from synthetic ABE mixture by CBV28014. Butanol (▲) adsorption from
fermentation broth as shown in Table 5.
Fig. 3 shows the comparison of extended Langmuir model predictions with the
experimental data from Tables 4 and 5. The highest deviations from the experimental
results occur at low concentrations and corresponding adsorption equilibria. The IAS model
results are shown in Fig. 4. The IAS model predictions are closer to the experimental data
than the simple extended Langmuir model. The maximum adsorption capacity for each of
the individual species on the zeolites differs, which can more adequately be handled by the
IAS model. The average difference between the adsorption predicted by both models and
experimental adsorption is 0.0032 and 0.0054 g g-1, respectively.
4.3.4. Qualitative description of adsorption equilibria*.
The octanol/water partitioning (see table 6) follows the trend butanol>butyric
acid>acetone>ethanol. This is mostly in line with the affinity following from the adsorption
isotherms (see figure 2). The Hildebrand and Hansen parameters show the affinity of the
specific compound for water to follow ethanol>butanol>acetone. Acetone and butanol
behave off key. Clearly not only compound and its interaction with water play a role, but
also the hydrophobicity of the sorbent. This is observed behavior, as e.g. [Ref 15]. Using a
Hildebrand/Hansen parameter will not provide a full picture and also hydrophobicity does
have a significant effect on the difference between water and solvent adsorption.
Table 6: Solubility data sorbates
Log Po/w [14]
Butyric acid
[Pa0.5] [14]
Table 7: Polarity scales for sorbents
Si/Al ratio
20 [ref 17]
23 [ref 16]
Function of Si/Al ratio. Specific data not available, b 1-silicalite
Index [18]
The hydrophobicity index for the used zeolites, see Table 7, are, if data is not directly
available for the specific zeolite, related to the Zsm-5 or Y-type structure with comparable
Si/Al content. The HI indices are approximately similar for CBV901 and CBV28014 as
both are high silica zeolites. The differences in adsorptive behavior are thus not only due to
direct surface interactions. The difference in structure plays a role [Ref 19]., namely a
micropore ZSM-5 type material (CBV28014) and micropore/mesopore containing Y-type
(CBV901), leading to an overall increase in water content. This difference in structure
changes the overall adsorptive properties of the silica crystals and results in the selectivity
of recovery for CBV28014 and CBV901 to be far more favorable for CBV28014, namely
274 to only 12.
4.4. Conclusions
CBV901 has the highest adsorption capacity for butanol of the three zeolites investigated,
but CBV28014 shows the highest affinity for butanol at aqueous butanol concentrations
below 2 g L-1. The competitive adsorption on CBV28014 showed a correlation with
hydrophobicity of the sorbate with the highest affinity for butanol > acetone > ethanol.
Single solute isotherms were successfully used for predicting multi-component adsorption.
The ideal adsorbed solution model predicts the competitive adsorption of butanol, acetone
and ethanol on CBV28014 more accurately than extended single site Langmuir model.
Adsorption from fermentative mixtures was also predicted successfully but butyric acid
needed to be taken into account as additional minor component.
4.5. Acknowledgments
The authors wish to thank Zeolyst International for providing the zeolite material, Ana
Lopéz-Contreras (Wageningen University and Research Centre) for providing the
fermentation broth and Max Zomerdijk for analytical support.
This project is financially supported by the Netherlands Ministry of Economic Affairs and
the B-Basic partner organizations ( through B-Basic, a public-private
NWOACTS programme (ACTS = Advanced Chemical Technologies for Sustainability).
4.6. References
[1] D. Antoni, V.V. Zverlov, W.H. Schwarz, Biofuels from microbes, Appl. Microbiol.
Biotechnol. 77 (2007) 23–35.
[2] P. Dürre,
acetone/butanol/isopropanol fermentation, Appl. Microbiol. Biotechnol. 49 (1998) 639–
[3] T.C. Ezeji, N. Qureshi, H.P. Blaschek, Butanol fermentation research: upstream and
downstream manipulations, Chem. Rec. 4 (2004) 305–314.
[4] N. Qureshi, S. Hughes, I.S. Maddox, M.A. Cotta, Energy-efficient recovery of butanol
from model solutions and fermentation broth by adsorption, Bioprocess Biosyst. Eng. 27
(2005) 215–222.
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Smith, Silicalite, A new hydrophobic crystalline silica molecular-sieve, Nature 271 (1978)
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liquors, Biotechnol. Lett. 4 (1982) 759–760.
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interfaces: unusually high affinity for hydrophobic, ultrastabilized zeolite Y, J. Phys. Chem.
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liquid solutions, AIChE J. 18 (1972) 761–768.
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Ge-ZSM-5 membranes, Micropor. Mesopor. Mater. 58 (2003) 137–154.
[13] M.T. Holtzapple, K.L. Flores, R.F. Brown, Recovery of volatile solutes from dilute
aqueous-solutions using immobilized silicalite, Sep. Technol. 4 (1994) 203–238.
[14] W.M. Haynes et al., Handbook of chemistry and physics, 92nd edition, internet version
2012, CRC Press.
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(2005) 318-331.
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zeolite beta, Micropor. Mesopor. Mat. 22 (1998) 1-8.
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by dealuminated zeolite Y, Micropor. Mesopor. Mat 99 (2007) 251–260.
4.7. Nomenclature
surface area of adsorbent (m2 kg-1)
liquid phase concentration (kg m-3)
single solute equilibrium liquid phase concentration (mol m-3)
Langmuir constant (m3 kg-1)
mass (kg)
mass equilibrium solid phase concentration (kg kg-1)
single solute equilibrium concentration in solid phase (mol kg-1)
maximum adsorption capacity (kg kg-1)
gas constant (J mol-1K-1)
temperature (K)
liquid volume (m3)
solvent free liquid phase mole fraction (–)
adsorbed phase mole fraction (–)
Greek symbol
spreading pressure (N m-1)
sorbate species i
all sorbate species
initial condition
Chapter 5: Desorption of butanol from zeolite material
Sorption-based recovery of 1-butanol from aqueous solution has been investigated focusing
on the recovery of butanol by desorption from the sorbate. Sorption isotherm,
thermogravimetric adsorption and differential scanning calorimetry experiments have been
used to determine the desorption behavior of butanol and water for two high-silica zeolite
adsorbents, CBV901 and CBV28014.
Carbon dioxide can be used as displacement agent for butanol recovery, with the butanolcarbon dioxide equilibria determining the carbon dioxide mass requirement for such a
For CBV901 desorption requires 2440 J per g of water and 1080 J per g of butanol. The
heat effects for CBV28014 are 2730 J/g (water) and 1160 J/g (butanol). A significant
difference in water content can be seen between both zeolite materials, with CBV28014
showing the least amount of water adsorption. The desorption rate of butanol from
CBV28014 is significantly slower than from CBV901.
A catalytic reaction, most probably dehydration, occurs around 200 oC during temperature
programmed desorption of butanol from CBV28014.
Published as: Desorption of butanol from zeolite material, A. Oudshoorn, L.A.M. van der
Wielen and A.J.J. Straathof, Biochem. Eng. J., 67 (2012), 167-172
5.1. Introduction
Microbial butanol production from renewable feedstocks has been receiving continuous
attention, because butanol has favorable fuel properties [1]. The microbial production of 1butanol will provide aqueous fermentation broth with approximately 10 to 20 g/L butanol
or total organic solvents [2,3,4]. Distillation is generally applied for product recovery, but
this is relatively energy-intensive [5,6]. Alternative recovery techniques are much less
developed. Frequently proposed alternative downstream processing methods are adsorptive,
extractive and pervaporation based recovery techniques. For adsorptive recovery of
butanol, high silica zeolites are selective and stable adsorbing agents [7,8,9,10]. Desorption
of butanol from the adsorbed zeolite phase is less frequently mentioned, but this is a crucial
process operation and needs to be taken into account when designing a complete recovery
Desorption of adsorbed species from a sorbent can be done by 1) pressure swing operation,
2) purge gas stripping, 3) displacement desorption, and 4) thermal swing operation [11].
Recovery by pressure swing operation is impractical, due to the very low vapor pressure of
butanol at ambient temperature. Recovery by purge gas stripping will require, also due to
the low vapor pressure, extreme volumes of purge gas. Therefore this paper focuses on
displacement desorption and thermal swing operation for desorption of butanol from high
silica zeolite.
Regeneration of an adsorbent using thermal swing operation is probably the most used
regeneration technique [11]. Thermal swing operations use the shift in adsorption behavior
as a function of temperature. Displacement of butanol from zeolite can occur if a newly
introduced displacement agent is adsorbed by the sorbent. After displacement, regeneration
of the sorbent phase by removal of the displacement agent is necessary if the sorbent is to
be reused. To allow relatively easy removal of the displacement agent from the sorbent, in
our opinion gaseous compounds at elevated pressures should be used. Fermentative
production of butanol will involve carbon dioxide [12], so this compound might be useful if
it shows sufficient adsorptive behavior.
Essential data necessary for the design of adsorptive recovery of butanol from fermentation
broth using high silica adsorbents will be determined.
5.2. Materials and Methods
5.2.1. Materials and sample preparations
1-Butanol (BuOH) was from Acros (purity 99.5 %). The powdered zeolites, CBV28014
(ZSM-5-Type) and CBV901 were from Zeolyst International, USA. All zeolites were
calcined at 600 oC for 8 h. For further details on the zeolites see table 1. The structure of
MFI and FAU crystals are described in literature [13,14].
Table 1: Manufacturer’s specifications of zeolite adsorbents and pore volume estimates.
Nominal cation
Surface area
Pore volume
5.2.2. CO2 and butanol competition
The adsorption of butanol on the investigated zeolites in the presence of dissolved carbon
dioxide was measured using a setup and experimental method similar to one previously
[15]. The aqueous butanol concentration was determined by gas
5.2.3. Gaseous butanol and water adsorption isotherm determination
The gas phase equilibria between butanol, water and zeolite (CBV901 and CBV28014)
were determined by gas phase pressure measurements in an enclosed temperature
controlled system (1,270 ml pressure vessel combined with a WIKA, EN 837-1
manometer). The loading of adsorbed species was determined by following the pressure
changes in the system at various amounts of sorbate introduced into the system. The liquid
phase butanol or water was injected into the system, in portions of approximately 0.1 to
0.15 gram per injection, up to a total amount of around 2.5 g. The amount of adsorbed
species was calculated from the difference between the observed pressure and the pressure
the total volatile species injected should have reached at that temperature if no adsorption
had occurred. In the loading calculation the ideal gas law was used to convert pressure into
mole of gas. Total volatile pressure in the system was always below the saturated vapor
pressure, to make sure no internal condensation could occur.
5.2.4. Differential scanning calorimetry (DSC)
Temperature programmed differential scanning calorimetry (Perkin-Elmer DSC-7) was
used to determine endothermic heat flow to zeolite and zeolite loaded samples during the
temperature change. This allowed the determination of the heat capacity of the zeolite
material and the heat uptake of adsorbed water or butanol when desorbing from the zeolite
Sample and reference containers were aluminum cups. For each DSC desorption run, the
heat uptake of an aluminum sample cup was measured and subtracted from the values of
the subsequent desorption run to obtain the net heat effect of the sample. The sample was
continuously flushed with nitrogen gas at 20 ml/min to remove desorbed components. The
DSC analyses were carried out using a programmed temperature profile, which applied a
heating rate of 10 oC/min, starting at 25 oC. By integration of the heat flow over time it is
possible to calculate heat effects per mass amount of sample. Integration of the DSC
endothermic heat flow allows calculation of the heat of desorption of the volatile species,
and also allows the calculation of the heat capacity of the zeolite materials. Similar heating
profiles have been used to determine desorption of volatile organics from micropore
material and activated carbon [16,17]. Initial isothermal steps of 5 min were programmed
to allow baseline correction. The samples of CBV901 and CBV28014 were loaded with
butanol, water or a mixture thereof by gas phase equilibration in a temperature controlled
vessel (25 oC) in the presence of liquid volatile species during 48 h. No difference in
adsorbed mass on the zeolite could be seen after 48 h equilibration compared to 72 and 168
h of equilibration. Liquids were pure water, pure butanol or 20 g/L butanol in water. An
amount of the loaded sorbent was then placed in an aluminum cup, after which a desorption
run was carried out. The mass of the sample with the cup was measured before and after the
run. The amount of sample material was usually between 10 and 15 mg total loaded zeolite
sample, including approximately 1 to 2 mg volatile component.
5.2.5. Thermogravimetric adsorption (TGA)
Thermogravimetric adsorption (Perkin-Elmer), was applied to sorbate-containing zeolite.
During the measurement the mass of a specific sample was monitored. The TGA analyses
were carried out using the same temperature program as in the DSC experiments (a heating
rate of 10 oC/min, starting at 25 oC). Also for the TGA experiments the samples of
CBV901 and CBV28014 (approximately 10 to 15 mg total sample mass) had been
equilibrated with butanol, water or a butanol-water mixture, following the same method as
described in the DSC section. The loaded sorbent was then placed in an aluminum cup,
after which a desorption run was carried out. TGA was performed at atmospheric pressure,
using nitrogen or air, 80 ml/min, as purge gas.
5.3. Results and discussion
5.3.1. Recovery by displacement
To be able to evaluate the potential of CO2 as butanol displacement agent, the isotherm of
butanol in the presence of carbon dioxide was measured for CBV28014 (figure 1). A
decrease of butanol adsorption on CBV28014 is observed with increasing CO2 pressure.
The difference between the data for 0 and 1 bar (CO2-saturated aqueous solutions
containing butanol) indicates that during adsorption the solute CO2 competes with butanol
and diminishes the total amount of butanol adsorption. Previously, this competition had not
yet been taken into account [9], but it can play a role especially as anaerobic fermentation
broths will usually be saturated by 1 bar CO2, corresponding to ~ 1.5 g/L in aqueous
solution. In figure 4, at 9 bar CO2 the aqueous butanol concentration is still relatively high,
indicating that this pressure is relatively low for using it to efficiently desorb butanol. Our
equipment did not allow higher pressures to be evaluated. Our results are in line with those
of [18], who preferred 1-butanol to CO2 for desorption of succinic acid from CBV28014.
Equilibria between 1-butanol and CO2 at elevated pressure show a low butanol content. For
CO2-based displacement at elevated pressure the CO2 the gas phase capacity for butanol
then determines to the economic effectiveness as it dictates the CO2 flow requirements.
Increase of pressure will allow larger butanol fractions at pressures over 60 bar [19], as will
an increase in temperature, at 333 K the fraction butanol is 3 to 4 times larger than at 303 K
[20]. As 1-butanol is only of the
possible C4 alcohols, it should be noted that for 2-
butanol the alcohol-CO2 gas phase equilibria are significantly more favorable, when
compared to 1- butanol and CO2, making this concept more feasible for this product.
Figure 1: Competitive adsorption of butanol from aqueous solution at 25 oC for CBV28014
under no (♦), 1 bar (□), and 9 bar (▲) CO2 -gas phase pressure. Lines have been added to
guide the eye.
5.3.2. Recovery by thermal swing
At 25oC, CBV901 can adsorb butanol up to 0.16 g/g and CBV28014 up to 0.12 g/g [9].
Adsorption isotherms of butanol from butanol vapor at elevated temperature are shown in
figure 2. At 110 oC the equilibrium loading of butanol is significantly lower on CBV901
than on CBV28014. The loading on CBV901 was reduced much further by increasing the
temperature to 150 oC. No data are given at 150 oC for CBV28014 because the observed
pressures were larger than expected on basis of the amount of butanol introduced in the
system. This suggests a catalytic reaction to occur producing additional gaseous molecules.
Figure 2: Butanol equilibrium loading on zeolite as function of butanol vapor pressure for
CBV28014 at 110 oC (♦) and for CBV901 at 110 oC (□) and 150 oC (▲). Lines have been
added to guide the eye.
5.3.3. Desorption rates
The desorption rates of butanol and a butanol and water mixture from CBV901 were
followed using TGA. Figures 3 and 4 are a function of temperature. The second abscissa
shows the time progression of the experiment. The observed sample mass has been
normalized to the amount of adsorbent present at the end of the experiment, assuming all
sorbate had been desorbed.
TGA desorption profiles normalized on final mass are shown in figure 3. The two top TGA
profiles show a significant amount of water to be present in comparison to pure butanol
adsorption. The large slope of the two top samples from 400 to 800 s indicates water is
removed at that stage. As all curves start behaving more similar after 800 s, it is assumed
that the remaining butanol is being removed in this stage.
Figure 3: Duplicates of TGA profile of the desorption of butanol (bottom lines) and
butanol/water mixture (top lines) from CBV901. Reading from top to bottom, the initial
mass of volatiles present in the samples were 6.91, 5.87, 2.50 and 1.55 mg. See Materials
and Methods for further sample preparation information.
Similar to the experiments on CBV901 the desorption of water, butanol and a mixture of
water and butanol from CBV28014 was determined. In figure 4 the normalized mass profile
can be seen. CBV28014 has a MFI channel structure and is very hydrophobic in nature.
This effect is seen by the additional water mass profile in figure 4. The overall water
content in the sample is very low, when the zeolite has been equilibrated with water. The
difference between butanol and butanol-water is thus also very small, as water is not
present in significant quantity, this opposed to CBV901. CBV901 showed significant water
uptake to occur [9]. Complete desorption of butanol from CBV28014 can be achieved at
temperatures below 200 oC. At ~ 200 oC a pronounced change in the slope of butanolcontaining samples can be seen, indicating a change in mechanism. This will be explained
in the later section when also looking at its heat effect.
Figure 4: Desorption profile of CBV28014 with three different loadings: butanol,
butanol/water and water. Reading from top to bottom the initial amount of volatiles present
were 2.5, 2.5 and 0.25 mg, respectively.
The mass transfer of butanol and water from the sorbent to the vapor phase was modeled
using a linear driving force relation, in order to obtain an estimate for the overall mass
transfer coefficient (ko) during the desorption runs:
 ko zeolite  (q  q* )
A relatively large flow of inert stripping gas was applied, so that the volatile species
concentration in the gas phase could be assumed to be close to zero. This implies the
equilibrium loading of the sorbate (q*) to be approximately zero, q* = 0. The current actual
loading of the zeolite (q) is seen as an average actual loading (q) of the zeolite. In general
for a gaseous system the overall mass transfer coefficient for a particle is determined by the
resistance in a gaseous film (external mass transfer resistance), the resistance in the internal
structure of the particle (pore diffusion), and the surface diffusion resistance of the sorbates
to the specific adsorption sites e.g. [21]. Due to the high temperature and purge gas flow the
external mass transfer resistance is probably not limiting. The observed overall mass
transfer is then mainly influenced by particle size and shape and by internal diffusion
through the pore structure of the zeolite. Structural differences between the zeolite particles
show the pore structure to be more constrained for MFI than for FAU types. Therefore a
stronger mass transfer limitation and lower mass transfer coefficient is expected for
Figure 5: The observed overall mass transfer coefficient of desorption of butanol from
CBV28014 (♦) and CBV901(□). Lines have been added to guide the eye.
The mass transfer rates, see figure 5, increase with temperature, as is expected and reaches
a maximum for CBV901 at around 95oC. The wide structure of the CBV901 in all
probability allows easy desorption of a large fraction of relatively loosely adsorbed butanol.
This is also later illustrated by the observed heat effects. The observed overall mass transfer
coefficient for desorption of butanol is lower for CBV28014 than for CBV901. The butanol
adsorption isotherm also showed a higher affinity for CBV28014 than for CBV901 [9].
This all shows the butanol on the CBV28014 to be stronger bound. For CBV901, after
reaching 90 oC, most of the sorbate has been removed and the low amount of butanol
present on the zeolite is limiting the overall mass transfer. The heat effects after 90 oC
behave in line with the decrease in sorbate removal rate, as can be seen later in figure 7,
implying no strong difference in binding energy or different adsorption sites play a role.
5.3.4. Heat capacities and heat of desorption
The DSC method monitors the endothermic heat flow (Hheating) in time necessary for the
system to follow the preset heating profile. The specific heat capacity (Cp) for a specific
non-volatile species can be calculated if the total measured heat effect is determined by the
DSC, as long as the sample mass (m) and heating range (Tend-T0) are known. The average
heat capacity of a zeolite sample is then:
C p , zeolite 
H heating
mzeolite  (Tend  T0 )
The heat capacities of CBV901 and CBV28014 as a function of temperature are shown in
figure 6. As can be seen from figure 3 the heat capacity of both zeolites increases slightly
with temperature and ranges for the listed temperatures from 0.85 to 0.98 J/(g.oC) for
CBV28014 and 0.95 to 1.1 J/(g.oC) for CBV901. This is in line with the reported heat
capacity of 1.0 J/(g.oC) for 1-silicalite [7]. The difference in heat capacity between the two
zeolites is expected as the zeolites structure differs strongly (ZSM-5 versus Y-Type). The
heat capacity is relatively low for the zeolites compared to water (the bulk of the original
aqueous feed phase). This is favourable with respect to the energy required for usage of
high silica in thermal desorption processes. Table 2 summarizes the heat properties of the
investigated components.
Figure 6: Heat capacity as function of temperature of CBV901 [Top line] and CBV28014
[Bottom line].
Table 2: Heat capacity and enthalpy of evaporation
N.A. d
CBV28014 0.95 b
95 d
N.A. d
(at 25 C), (at 75 C), Calculated from (Tb-T0).Cp, with
This paper
This paper
T0 = 25 oC, d(per 100 oC). : Not
applicable, eat boiling temperature Tb.
For validation of the DSC method the heat of evaporation of pure water was measured in
the same setup, using the applied temperature profile. The measured heat was 2570 J/g.
This corresponds to the sum of the sensible heat of water between 25 and 100 oC and the
evaporation enthalpy at its boiling point, according to the data in Table 2.
Similar to the heat capacity the heat of desorption was monitored using DSC. Integration of
the DSC signal allows calculation of the overall desorption heat. The heat uptake for
butanol and water are shown in figures 7 and 8 for CBV901 and CBV28014, respectively
as a function of temperature. The second abscissa shows the time progression of the
Figure 7: Heat of desorption of water and of butanol from CBV901 (0.50 mg and 2.24 mg,
Figure 8: Heat of desorption of water and of butanol from CBV28014 (0.70 mg and 2.14
mg, respectively).
Figure 7 and 8 show water to desorb predominantly at low temperatures compared to
butanol. The relative high amount of water present in CBV901 is desorbed easily and the
water is thus not strongly bound, which results in the peak in the desorption profile.
However, figure 7 does show butanol to desorb over the entire temperature range, which is
in line with the TGA profile in figure 3. The difference in desorption heat profiles between
water and butanol is more significant for CBV28014. This indicates a significant difference
in butanol adsorption to exist between both zeolite structures. This large difference in
desorption temperature between both components can be used to perform temperature
programmed desorption, in which almost pure fractions of water and butanol are obtained,
which has already been mentioned for 1-silicalite [23].
The heat required for the desorption of volatile components from the investigated zeolites is
close to the sum of heat of evaporation and sensible heat of the volatile species
investigated. This means the heat of adsorption is small when compared to the heat of
evaporation. Figure 4 shows the desorption rate to change significantly around 200 oC. The
heat effect of the desorption in figure 8 also shows a change in behavior to occur here. The
desorption of butanol from CBV28014 shows an endothermic peak at around 200 oC. This
endothermic behavior can occur from a chemical reaction occurring. The heat effect seen in
figure 8 is probably due to a dehydration reaction, leading to butene formation. This
reaction has been observed using granulated zeolites of CBV28014 [24].
In table 3 the total heat of desorption of desorbed volatile species is shown. The overall heat
of desorption is shown per gram of desorbed component, under saturated loading.
Table 3: Heat effect of desorption per amount of BuOH and water.
Heat uptake
The heat of desorption is less pronounced for butanol than for water. This is not unexpected
when comparing their evaporation enthalpies (Table 2). The heat of desorption for butanol
is 275 J/g and 355 J/g above the heat of evaporation of the pure species, for CBV901 and
CBV28014, respectively. The heat of desorption of water shows a ~5 % deviation from the
heat of evaporation of the pure species. The differences are -130 and 160 J/g for CBV901
and CBV28014, respectively. This inversely indicates that the adsorption heat at 25oC will
be relatively high for butanol as compared to water, which is a consequence of the stronger
adsorption of butanol. The measured heat effects for CBV28014 also indicate stronger
molecular interactions between the volatiles and this zeolite, compared to CBV901. The
stronger hydrophobic interactions between CBV28014 and butanol increase the difficulty
of thermal desorption.
5.4. Conclusions
The adsorption isotherm of aqueous butanol on CBV28014 changes significantly if the
aqueous solution is at equilibrium with up to 9 bar of gaseous carbon dioxide. This
difference can be exploited in a pressure swing desorption operation.
The small competitive adsorptive behaviour of CO2 relative to butanol should be taken into
account when adsorbing butanol from fermentation broth saturated by 1 bar of carbon
The heat capacities of CBV901 and CBV28014 are (at 75 oC) 0.95 and 0.85 J/(g.oC),
respectively. These relative low heat capacities pose no direct limitation on usage of high
silica in thermal swing desorption processes.
Desorption of butanol from CBV901 and CBV28014 by thermal swing operation shows the
overall desorption heat effect for water to be marginally affected. Butanol is bound stronger
by CBV28014 than by CBV901. The heat of desorption for butanol is 275 J/g and 355 J/g
above the heat of evaporation of the pure species, for CBV901 and CBV28014,
respectively. Also, the desorption of butanol from CBV28014 shows a catalytic reaction to
occur around 200 oC.
The desorption rate of butanol from CBV28014 is significantly slower than from CBV901.
The increased mass transfer resistance can be caused by a difference in pore diffusion
characteristics due to the smaller pore channels of CBV2014. Complete regeneration of
both zeolites is possible using thermal operation.
5.5. Acknowledgements
The authors wish to thank Zeolyst International for providing the zeolite material and Ben
Norder of the Delft University of Technology, Chemical Engineering group, for technical
support. This project is financially supported by the Netherlands Ministry of Economic
Affairs and the B-Basic partner organizations ( through B-Basic, a publicprivate NWO-ACTS programme (ACTS = Advanced Chemical Technologies for
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O. Levenspiel, Chemical reaction engineering, (1999), John Wiley and Sons.
W.M. Haynes et al., Handbook of chemistry and physics, 92nd edition, internet
version 2012, CRC Press.
N. B. Milestone, D. M. Bibby, Concentration of Alcohols by Adsorption on
Silicalite, J.Chem.Technol.Biotechnol. 31 (1981) 732-736.
V. Saravanan, D. A. Waijers, M. Ziari, M. A. Noordermeer, Recovery of 1-butanol
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Chapter 6: Short-cut calculations for integrated product recovery options in
fermentative production of bio-bulk chemicals
Micro organisms are generally sensitive to high concentrations of products that they
excrete. Such product inhibition and toxicity effects can significantly be reduced by the
integration of fermentations with separation technologies to remove the products
continuously. Cost-calculations are required to select the preferred integration method. This
paper presents a shortcut calculation method that provides easy interpretable results to
assist the reader in making rational design choices. The method distinguishes four main
cost categories, being capital and operational expenditure for both fermentation and product
recovery. We have applied these cost correlations to the production of three typical biobased bulk chemicals. The results show the origin of the most significant costs in the
investigated integrated bio-processes. The presented method can be used to direct future
research efforts and might assist in evaluating the impact of integrated product recovery
techniques on total production costs of bio-chemicals.
Keywords: Integrated product recovery, Lactic acid, Butanol, Phenol, Review, Cost
calculations, Fermentation, Separation technology
Published as: Short-cut calculations for integrated product recovery options in fermentative
production of bio-bulk chemicals, A. Oudshoorn*, C. van den Berg*, C.P.M. Roelands,
A.J.J. Straathof, L.A.M. van der Wielen, Process Biochemistry 45 (2010) 1605–1615
*Co-Authors, both authors contributed equally
6.1. Introduction
Fermentative conversions of substrates into desired chemical products have been
commercialized over time by pharmaceutical and chemical industry. Biocatalysts usually
show a high selectivity towards the product and therefore biological conversions can be
used in a wide range of fields. Currently, there is a focus on sustainable resources and
production systems. This paper deals with the production of bulk chemicals by de novo
fermentative conversions with integrated product recovery. Microbial production of biobased chemicals is often hampered by product inhibition or product toxicity [1]. The
performance of product inhibited fermentations can be enhanced by removal of the product
during fermentation, usually referred to as integrated product recovery [2]. Integrated
product recovery is not just a downstream operation, but is directly influencing the
performance of the fermentor. Thus, a direct coupling exists between the performance of
the fermentation and product recovery (see Fig. 1). The system shown in Fig. 1 can be seen
as an integrated production system. The (economic) performance of the integrated
production system is an optimization problem involving multiple unit operations. Table 1
shows some available integrated product recovery operations for bio-chemicals. This paper
details the integrated microbial production of lactic acid, 1-butanol and phenol. The main
goal of this work is to show the economic effect of integrated product recovery on
production systems. The results only interpret the effect of product recovery. Other
possibilities for stream integration, e.g. substrate utilization and waste management, are not
taken into consideration. The obtained economic outline can be used to further focus
research and development efforts of integrated product recovery designs.
Fig. 1. Fermentation with integrated product recovery.
Table 1 Some possible recovery operations for bio-based bulk chemicals from fermentation
New Phase
composition change
Composition change
Composition change
Composition change
Composition change
Pervaporation [3–6]
Gas stripping [7–10]
Extraction [11–15]
Pertraction [16–20]
Adsorption [21–26]
6.2. Product characteristics
6.2.1. Fermentation
Fermentations can be quantified by product yield, volumespecific productivity and critical
concentration. Table 2 lists the fermentation performance data that are used to define the
specific fermentations in this paper. Extensive literature on the production of these
chemicals is available. Lactic acid shows the highest critical concentration, volume-specific
productivity and yield. The butanol case can be characterized by a moderate volumespecific productivity, product yield and critical concentration. Phenol is the most inhibiting
of the three compounds, with the lowest critical concentration, volume-specific productivity
and yield.
6.2.2. Lactic acid
Lactic acid, when compared to butanol and phenol, has the highest aqueous solubility and a
negligible vapor pressure. The negligible vapor pressure renders recovery options involving
an auxiliary gas phase (such as pervaporation) non-feasible. Therefore, only recovery by
adsorption, extraction and pertraction are evaluated. Adsorptive based recovery involves
adsorbents, which can range from (organic) resins to inorganic materials, like zeolites
[35,36]. Only lactic acid adsorption on zeolite is considered in this paper, because the high
stability of the zeolite adsorbent material is assumed to be favorable.
Solvent selection [37–39] is crucial in the design of extractive recovery systems. Among
the best extractants for L–L extraction/pertraction of lactic acid is tri-n-octylamine, which
has a lactic acid partition coefficient of 4 and dissolves a low amount of water (0.51
mass%). Regeneration of the extractant will be performed using a base/acid wash.
Overall mass transfer coefficients for pertraction based recovery have been reported by
Huang et al. [40]. Desorption and extractant phase regeneration operations are also
constrained by the negligible vapor pressure. Desorption needs a chemical auxiliary,
introduced in the form of a base or solvent wash. For the adsorption case, desorption using
an ethanol washing step is assumed[41].
The energy demand during regeneration is then determined by ethanol evaporation, see also
Section 6.4.4.
6.2.3. Butanol
Due to the moderate aqueous solubility of butanol a wide range of possible product
recovery techniques is available [42,43]. Furthermore, butanol has the highest vapor
pressure of the three cases, explaining the relatively large amount of literature concerning
gas-phase product recovery [44].
Table 2 Product characteristics for evaluating fermentation.
Maximum volume
specific productivity yield
Lactic acid
[27], [28], [29], [30],. [31]., [32]., [33]., [34].
Numerous adsorptive recovery operations, with various adsorption materials, can be used
[43,45]. Often zeolite material has been suggested, and this can be used in a granulated
form, allowing easier handling of the material due to increased particle sizes. However,
additional water will be adsorbed due to the presence of macropores. Macropores can
make-up 40% of the material [46].
The extractant considered for L–L extraction/pertraction of butanol is Cyanex-923, which is
a liquid phosphine oxide with a butanol partition coefficient of 14.8, which is the highest
butanol partition coefficient recorded to our knowledge. Cyanex-923 dissolves 6 mass% of
water and regeneration will be performed using a heat strip [47]. Butanol can be
regenerated from the adsorbent or extractant phase using an evaporation procedure. Butanol
removal from fermentation broths in current industrial practice is performed using
distillation but this is not an integrated process.
6.2.4. Phenol
Phenol removal from aqueous phases is discussed in [48]. Due to the severe toxicity of
phenol, only low product concentrations are in the aqueous fermentative phase. The low
concentrations, in combination with the relative low vapor pressure of phenol [49], will
make pervaporation based recovery from a overall process perspective non-feasible.
Adsorption of phenol is possible [50–53]. Activated carbon, which was reported by Costa
and Rodriguez [54] will be used here for adsorption based calculations. Costa and
Rodriguez [54] reported an expression for loading time for multiple experimental runs of a
fixed bed, depending on liquid flow-rate through the column. Regeneration of the sorbate
will be performed using ethanol. For L–L extraction/pertraction, again, an organic
extractant such as Cyanex-923 will probably work best and is used in this paper due to
phenol’s partition coefficient of 1100 coupled to a modest water uptake of 6mass%. A base
wash with subsequent acid neutralization will be used for both extractive phase
6.3. Methods: definitions, assumptions and boundaries
The integrated system, shown in Fig. 1, can be described in various degrees of detail. The
system is defined as consisting of two coupled unit operations, the fermentation and the
integrated product recovery. The main goal is to show the general economic behavior of
such an integrated system. Therefore, the economic performance has been divided into two
main costs, the capital expenditure (CAPEX) and operational expenditure (OPEX). Both
fermentation and product recovery influence the CAPEX and OPEX. This paper reduces
the complexity of the optimization problem of integrated systems, by linking all four cost
factors to one main process variable, the product concentration in the fermentor, as will be
shown in the next section. General assumptions, simplifying the analyses, are used in this
paper: the production amount is 100,000 tons annually; the fermentation and integrated
product recovery are continuous operations; the microorganisms are retained in the
fermentation unit and do not grow significantly; an annual production time of 8000 h is
assumed for sizing of the fermentation equipment. Furthermore, literature values of product
fluxes, loading capacities, volume-specific productivities and carbon yields have been
obtained to allow modeling of the behavior of the unit operations. A mass balance of the
product over the unit operations, shown in Fig. 1, shows the system to operate in a steady
state as long the amount produced in the fermentation equals the amount of product
removed during product recovery. This assumption is valid as long as a possible bleed or
purge is neglected. The productivity of the fermentation (Qfer) and the product flow (Frec)
are then equal to:
Qfer = Frec (kg/h)
Cost optimizations of the system shown in Fig. 1 are thus possible for any production
amount, if both Qfer and Frec are used to size the process equipment. The specific method
applied to calculate the unit operations are detailed per recovery operation for each product
and can be found in Section 6.4, while using the details given in Section 6.2.
6.3.1. CAPEX fermentation
The capital expenditure required for the production of biochemicals depends on the size of
the fermentation vessel. In order to evaluate the annual capital costs on the fermentation
side for a required productivity (kg/h), estimations should be made on the volume-specific
productivity of the fermentor ((kg/m3)/h).
6.3.2. OPEX fermentation
In the fermentation, the feedstock is converted into a bio-chemical. The dominant
operational expenditure during fermentation is assumed to be determined by the feedstock
cost. The OPEX of the fermentation can thus be expressed as a (cost) function of the
product yield on substrate depending on the product concentration in the fermentation.
6.3.3. CAPEX product recovery
The capital expenditure is determined by the individual costs of the product recovery unit
operations. Most mass transfer operations can be expressed per area of transfer equipment.
Then the costs depends on the product flux of recovery option. The flux can be calculated
using equations presented in Section 6.4.
6.3.4. OPEX product recovery
The operational expenditure during the product recovery is determined by the energy costs
and the costs of any auxiliary material used (such as acid/base). The energy requirement, as
function of the product concentration entering the capture step, is needed for the calculation
of the OPEX. The energy requirement during recovery is dependent on the amount of
product and auxiliary phase, and the process conditions. The auxiliary phase is the phase
facilitating separation of the product compound from the aqueous feed phase. Key
parameter is the selectivity of recovery as it details the mass amounts involved in the
recovery operation [43].
6.3.5. Overview costs functions
Combining the previous Sections 6.3.1–6.3.4, it can be generally stated that main cost
considerations in the production of biochemicals arise from the four parameters given in
Table 3. In Section 6.4 we will detail the modeling and costs per operation.
Table 3 Cost determining parameters for fermentation and product recovery.
Product yield on substrate
Product recovery
Product mass flow
6.4. Modeling
The modeling of the fermentation and product recovery requires background process data.
First, the process specific data are listed, followed by the method for cost calculation. All
cost correlations presented in the next sections can be replaced by readers’ personal
correlations. Therefore, these relations should be used as a general guideline for making
cost estimations. In this paper we will discuss the three different cases (phenol, 1-butanol
and lactic acid) where the bio-chemicals vary from severely toxic to relatively non-toxic
towards the host micro-organism. Furthermore, these products have varying yields on
substrate and varying physical properties such as vapor pressure, aqueous solubility and
extractability, amongst others.
6.4.1. Volume-specific productivity of fermentation
In order to estimate CAPEX on the fermentation side, volumespecific productivities should
be known. However, inhibition kinetics and final product titers are difficult to predict and
are organism- and product-specific. When products are removed from the microbes’
environment, the yields on substrate and/or volumespecific productivities can increase
significantly. Fig. 2 shows a correlation between the aqueous solubility of compounds and
the concentration inhibiting the microorganisms producing them, adapted from Straathof
[55]. The so-called critical concentration will be set as the final product titer.
Fig. 2 shows critical concentrations for a wide range of products from biological
conversions. Critical concentrations are limiting product concentrations in the fermentation
unit and are therefore also the final attainable concentration when no product removal is
applied. Estimating a priori a volume-specific productivity Qmax is not possible since this
depends on whether the fermentation product is a primary or secondary metabolite and also
on the micro-organism. Therefore, maximum volume-specific productivities found in
literature were used as Qmax (see Table 2). To estimate the productivities, Eq. (2) from
Luong et al. [36] is used, which is a simplified representation of real product inhibition
Q  Qmax 1  ( ferm )β 
Ccrit 
Fig. 2. Least squares fit between solubility and critical concentration of biological
conversion products. The vertical stripe containing ellipsoid contains severely inhibiting
compounds such as alkaloids/pharmaceuticals; the squares containing ellipsoid represents
highly inhibiting molecules which are mostly fine-chemicals; the ellipse with horizontal
lines contains relatively non-inhibiting molecules and mainly consists of typical bulk
Fig. 3. Assumed influence of product concentration in fermentation broth on
volume-specific productivity and yield.
The β is an inhibition constant, which is set to 1.69, like in [56]. The volume-specific
productivity provides the relevant process information necessary to estimate the size of the
fermentation equipment. The listed volumetric productivities in Table 2 depend on the cell
density present. High cell density fermentations can show higher volume-specific
productivity. The high cell density fermentation will require the negation of an existing
limitation, e.g. mass transfer limitation.
6.4.2. Fermentation yield
Product yields on glucose (Yp/s) are directly coupled to the concentration of the product in
the aqueous phase in a similar way as the volume-specific productivity equation. Products
such as phenol, which are not produced by the microbial catabolism, will have product
yields on glucose approaching zero near the critical concentration due to their molecular
toxicity. Although product yields depend on type of microbial host and metabolic route, a
simple equation is required. Product yield can be calculated using Eq. (3), which is
analogous to Eq. (2), but it is outside the scope of this article to support it with data.
 C
Yp/s =Yp/s max 1-( ferm )β 
 Ccrit 
In Fig. 3 the yields and volume-specific productivity are plotted as a function of the
fermentation broth product concentration. It should be noted that yields of catabolic
products are set to a constant of 90% of their maximum theoretical yield since they are
produced by the maintenance reaction. The operational costs on the fermentation side are
dependent on substrate costs. However, renewable raw materials experiences large price
fluctuations, and prices of 2nd generation feedstock derived materials, such as
lignocellulosic hydrolysate are not yet known. Therefore, this paper defines one C-6
feedstock, glucose, valued at 0.25 euro/kg, which is done for transparency reasons.
Additional calculations using actual feedstock and feedstock prices can optionally be
carried out and compared the results shown here. As the glucose costs are set, for the
production of 100,000 tons/year the operational expenditure is a function of the product
yield on glucose and is shown in Fig. 4.
6.4.3. Product recovery fluxes
The capital expenditure for recovery operations depends strongly on the type of recovery
applied and product concentrations in the feed and product recovery phase. The cost of the
recovery equipment is directly linked to the costs of the equipment.
Fig. 4. OPEX of fermentation as function of the product yield on glucose.
Sizing the recovery operation will be based on the mass transfer of the chemicals involved.
The mass transfer is a function of the product concentration and will be calculated using
Eq. (4).Most process operations require regeneration of the auxiliary phase used to capture
the product. Using Eq. (4), the total interfacial area will be calculated assuming that the
concentration in the product recovery phase remains zero.
Frec = K ov  A eff   cferm  p 
For the listed recovery techniques, the relations linking the product concentration, via the
equipment size, to the expected capital expenditure, are given in the next section and are
obtained by comparing scientific literature data. In the next sections different recovery
techniques are discussed in terms of initial fluxes and costs associated with the operation. If
the same unit is used in both operations, either the capturing step or the recovery step
determines the size of the equipment. The unit operations are primarily sized in this paper
based on the capturing operation. The process times for any regeneration operation is
assumed equal the capturing step if no data is readily available.
119 Adsorption
The adsorption and desorption of the desired product are necessary operations during
adsorptive recovery of a bio-chemical. Most adsorption processes are column operations
and overall mass transfer coefficients can be derived from numerous correlations and
approximations. Transport rates for adsorption process for spherical particles will be
approximated by [57]:
k  15 
Rp 2
Eq. (5) is valid for pore diffusion controlled adsorption. The diffusion coefficient (Daq) can
be calculated using the Wilke–Chang correlation [58].
K ov  k 
Eq. (6) describes overall mass transfer for spherical particles, where pore diffusion is the
main rate-determining step. However, it is not the goal of this paper to fully calculate the
adsorption and desorption profiles. When Kov, the amount of relative interfacial area and
feed concentration are known, initial fluxes can be calculated using Eq. (4). Pervaporation
Pervaporation is carried out with aid of membrane material. A characteristic parameter is
the trans-membrane flux. Generally, the flux is expressed as a function of the retentate
product concentration and membrane thickness. Literature data expressing the membrane
flux can be found for various systems. No data on pervaporation, at fermentation
temperatures, were found for lactic acid and phenol since these compounds have a low
vapor pressure. Therefore only butanol will be discussed in this paper [59,60]. Transmembrane flux is mainly dependent on membrane material, temperature and membrane
thickness. Many different membrane materials can be used and literature is extensive. An
overview of membranes and their performance is given by Vane [3]. When product flux
and production amounts are known the required membrane area can be calculated. The
CAPEX is calculated from membrane and membrane housing costs. Then, capital
expenditure can be estimated using membrane prices and lifespans such as in Appendix B.
120 L–L extraction
Estimation of CAPEX of an in-stream L–L mixer-settler extraction step requires values of
the product fluxes into the organic phase. To calculate product fluxes, one needs to know
the required amount of interfacial contact area, and molecular diffusivities in aqueous and
organic phase. The required amount of interfacial area needed depends on the speed of
agitation in the mixer-settler. Paulo et al. [61] found an average Sauter mean droplet
diameter over a wide range of Reynolds numbers. Therefore, we apply an average drop
diameter (dvs) of 3.4mm in our calculations. It is assumed that the droplets will have an
aqueous film layer (Δz) and mass transfer coefficients in this layer will be calculated using
Eq. (7).
K aq 
These droplets are assumed to be rigid, resulting in a Sherwood number of 6.6. Organic
phase molecular diffusivities can be calculated using the Minhas-Hayduk correlation [62].
When Nsh, Dorg and dvs are known, the Korg value can be calculated using:
(Nsh )ov 
K org  dvs
 6.6
The overall mass transfer coefficient is calculated using Eq. (9):
K ov K org  p K aq
When overall mass transfer coefficients are known, the minimum amount of effective
interfacial area can be calculated using Eq. (4), assuming the concentration in the loading
phase is zero upon contact with the fermentation broth. To size mixer-settlers, the
interfacial area per volume of mixer-settler can be calculated using [63]:
a 
6  d
d vs
Фd is the solvent phase hold-up and will be set to 0.3.Now sizing and cost-calculations can
be performed using cost equations, and can be found in Appendix B. Pertraction
To estimate CAPEX on hollow-fiber pertraction units, mass transfer coefficients and
subsequent product fluxes need to be estimated. High organic/water phase partition
coefficients result in mass transfer limitations at the aqueous side [64]. Therefore, it is
assumed extraction rate is determined by the resistance at the aqueous side [65]. Therefore,
flux calculations involving a pertraction unit will be simplified using Eq. (11).
K ov K aq
Kaq will be calculated using Sherwood numbers, and has been normalized
for 1m length-scale:
N sh 
K aq  di
 di 2  aq
 1.62 
 D
1/ 3
Table 4 Phase transition properties of pure bio-based bulk chemicals and water [66].
Tm [◦C]
Tb [◦C]
Lactic acid
However, if products have poor organic/water phase partitioning behavior, mass transfer is
influenced by membrane and organic phase as well, making it more complex. For these
cases, overall mass transfer coefficients are taken from literature. After obtaining overall
mass transfer coefficients, product fluxes can be calculated using Eq. (4). Subsequently,
sizing of pertraction can be performed.
6.4.4. Product recovery
The regeneration energy penalty of the product recovery can be expressed by the selectivity
of the capturing operation. This selectivity is defined as the ratio of the product
concentration and the water concentration in the product recovery stream to that of the feed
stream (see Eq. (13)).
[Cproduct /Cferm ]perm
[Cproduct /Cferm ]feed
Energy input in the product recovery is one of the key parameters when describing its
economic feasibility. Based on a steady state flow process, the mass-specific enthalpy flow
is given by Eq. (14).
H  Qh  W
ΔH is the mass-specific enthalpy change, Qh is the amount of energy added to the system
per mass (kJ/kg) and W is mass-specific work (kJ/kg). The process enthalpy of both feed
phase and product phase needs to be calculated in order to estimate the energy needed for
the recovery operation. This means the mass flows of the species need to be coupled to the
enthalpy balance. The nonideality of the product phase in comparison to the feed phase
occurs is not considered here. However, the values of Qh and W applied on and by the
system are dependent on the technique used. One important assumption here is that an
azeotrope will not be formed during the regeneration of the product recovery phase,
resulting in underestimated energy costs. Usually, the product capture by integrated
recovery brings the product from a dilute aqueous stream to an enriched stream containing
concentrations of up to 50 mass% product. When starting with 20 g/l of product this means
that 49 out of 50 water molecules are removed in the concentration step (relative to a
product molecule). Most of the energy requirements are in the product recovery step.
Upgrading diluted aqueous streams up to 50 mass% of product will consume the majority
of energy required for product recovery. Subsequent product purifications steps will require
significantly less energy input (Table 4). As mentioned the unit operations are primarily
sized in this paper based on the capturing operation. The process times and thus required
capital expenditure for any regeneration operation is assumed equal the capturing step if no
data is readily available. Adsorption
Desorption of product from the adsorbent can be performed by either heat swing or solvent
wash. Organic solvents, e.g. ethanol, are used when products with a low vapor pressure are
involved. The loaded solvent will still need to be regenerated using temperature swing. The
amount of energy input required, is determined by the solute concentration in the washing
solvent phase. It is assumed that the total column volume minus adsorbent backbone
volume will be filled with washing solvent. All of the product will be desorbed by the
solvent and energy-efficiency calculations can be performed using Eqs. (13) and (14).
For the butanol case, being a typical example of a high vapor pressure product, a heat strip
is assumed. Pervaporation
The pervaporation of a product requires the feed phase to partly evaporate and form a
separate vapor phase. This can be achieved either by heating the feed or by forming a low
pressure (vacuum) at the vapor side of the membrane. The low pressure can be maintained
by either pumping or cooling the product side of the
pervaporation unit. Our method calculates the energy involved in the main phase transition
of the product phase. The product phase is usually brought from liquid to vapor and back to
liquid phase. The total energy requirement is calculated by applying Eq. (14) to the known
separation selectivity of the pervaporation operation and the chemical properties shown in
Table 3. Butanol is the only case where pervaporation is calculated and selectivity and
fluxes are taken from Hickey et al. [67]. Extraction and pertraction
The main OPEX costs for extraction and pertraction consist of solvent purchase and
regeneration costs. The solvent purchase costs depend on type of solvent. Removal of
hydrophilic products needs so-called reactive solvents (such as tertiary amines or phosphine
oxides) to reach high product concentrations in the organic phase. However, these reactive
solvents also dissolve significant amounts of water, resulting in an expensive regeneration.
It was stated in Section 6.4.4 that the selectivity of the organic phase towards the product
will determine the regeneration costs. The selectivity can be calculated when the water
solubility into the organic phase is known, together with the product partition coefficient.
Solvent losses due to their aqueous solubility do not occur in the simplified process scheme
of Fig. 1. The amount of solvent needed can be calculated from the total amount of
interfacial area needed is known, which can be calculated using Eq. (10).
6.4.5. Costs
All the individual chemical and process equipment prices are compiled in Appendices B
and C. Total manufacturing cost (TMC) is calculated by adding the OPEX and CAPEX of
both the fermentation and capturing operation. A factor of 4.93 is used to convert purchased
equipment costs (PCE) into fixed capital, allowing calculation of the CAPEX of the two
defined unit operations [68]. The reader might replace this CAPEX/OPEX calculation by
other reasonable ones [69,70]. A linear depreciation over 12 years was applied when
calculating cost functions. The resulting cost functions taken for the individual operations
are detailed in Appendix B.
6.5. Results and discussion
The subsequent figures show the cost lines that predict the capital expenditure and the
operational expenditure for both fermentation and recovery operation as a function of
product concentration in the fermentation broth.
6.5.1. Lactic acid
In Figs. 5–7 a comparison is shown of L–L extraction, pertraction and adsorption for
product recovery techniques. It can be observed that extraction has the lowest total
manufacturing cost (TMC). The low TMC is a result of the relatively low CAPEX costs of
the mixer-settlers units and relatively low regeneration costs. Pertraction has a higher
CAPEX demand as it includes large membrane area costs. The CAPEX demand of the
membrane operation is close to the TMC and the other cost contributions are almost on the
base-line in Fig. 7. Extraction and pertraction require a base-wash and acid neutralization
step costing about ~20% of the TMC. Running the fermentation at concentrations lower
than ~18 kg/m3 shows a steep increase in TMC for adsorption and pertraction. The
concentration of lactic acid in the adsorbent material will be low, resulting in low
concentrations in the washing solvent and subsequently low regeneration efficiency.
Similarly, the pertraction operation is strongly affected by operating at low product
concentrations. An increase in product concentration will increase the costs of the
fermentation. This can be seen most clearly in the TMC profile for the extraction system in
Fig. 6. Also, looking only at operational costs, TMC is affected strongly by glucose price.
Doubling the price of glucose would increase the TMC by approximately 43% of the
integrated recovery using extraction. The caption to be used in Figs. 5–14 can be found in
Fig. 5.
Fig. 5. Recovery of lactic acid by adsorption as function of the concentration in the
Fig. 6. Recovery of lactic acid by pertraction as function of the concentration in the
Fig. 7. Recovery of lactic acid by extraction as function of the concentration in the
6.5.2. Butanol
A comparison of adsorption, pervaporation, L–L extraction and pertraction for product
recovery is shown in Figs. 8–11. Due to the product inhibition of butanol, lower volume127
specific productivities require larger fermentation vessel volumes, resulting in higher
CAPEX. The fermentative production of butanol using integrated recovery by
pervaporation has an overall minimum, when operating the fermentation at approximately
14.8 kg/m3 butanol. At lower butanol concentrations the product flux through the
membrane is relatively low. The low product flux leads to an increase in membrane area
and the relatively low butanol concentration in feed phase results in a subsequently high
energy demand for regeneration. The TMC increases strongly at low butanol
concentrations. The butanol trans-membrane flux remains mostly below 1 kg/(m2 h). An
increase in butanol concentration will negatively influence the fermentation productivity,
shifting more costs towards the fermentative operation. The lowest TMC is predicted for an
adsorptive based recovery at 3.33 kg/m3 butanol in the fermentor. The amount of water
adsorbed relative to butanol, dictates the energy efficiency of the overall process and
energy costs dramatically increase when more water is adsorbed. The operational cost of
this integrated product removal dominates the TMC for extraction. The result is a relatively
high price even at low fermentation concentrations. It should be noted that the lowest TMC
is at a higher butanol concentration for extraction than for adsorption. The predicted TMC
values of the butanol removal techniques area relatively close to each other, explaining the
amount and diversity of research that has been performed on the subject of butanol
recovery [43].
Fig. 8. Recovery of butanol by adsorption as function of the concentration in the fermentor.
Fig. 9. Recovery of butanol by pervaporation as function of the concentration in the
Fig. 10. Recovery of butanol by pertraction as function of the concentration in the
Fig. 11. Recovery of butanol by extraction as function of the concentration in the
6.5.3. Phenol
Adsorption, pertraction and extraction recovery of phenol are shown in Figs. 12–14,
respectively. The microbial production of phenol does not allow concentrations in the
aqueous phase above 0.47 kg/m3. An efficient system needs to have a high affinity of the
auxiliary phase for phenol. All fermentation costs show a steep increase in TMC when
product toxicity in the fermentation becomes the dominant factor since yields on glucose
decrease. For adsorption, the OPEX of this integrated recovery is the main cost factor. This
implies that the phenol concentration and the corresponding adsorption equilibrium are too
low for efficient desorption of the adsorbent material with ethanol. Pertraction has a lower
TMC for the recovery of phenol concentrations. However, the best technique for integrated
processing seems to be extraction since this technique lacks the membrane investment
compared to pertraction. TMC for phenol production with extractive recovery is relatively
low for the phenol concentration range 0.05–0.28 kg/m3. Due to the relatively
volumetric productivity during fermentation and the high selectivity of extraction the
fermentation costs dominate the costs profile of extractive recovery. When phenol
productivity significantly decreases, the TMC increases due to higher fermentation
Fig. 12. Recovery of phenol by adsorption as function of the concentration in the fermentor.
Fig. 13. Recovery of phenol by pertraction as function of the concentration in the
Fig. 14. Recovery of phenol by extraction as function of the concentration in the fermentor.
6.6. Conclusions
Cost predictions for the biological production of lactic acid, butanol and phenol, using
integrated product recovery techniques, have been constructed. The cost predictions are a
result of the application of the shortcut methods derived in this paper, as shown in Sections
6.3 and 6.4.We believe the accuracy of the methods to be such that the behavior of cost
trends in Section 6.5 is correct. However, the absolute cost values are still rough
approximations. Therefore, the shortcut method may be used to quickly select the process
option to be developed further and to be analyzed in much more detail according to
conventional methods. The three chemicals show a wide range of physical properties and
fermentation parameters resulting in different cost contributions. Whenever the product
recovery operation is efficient at low concentrations or inexpensive, the production by
fermentation can take place at lower product concentration. Both the operational and capital
expenditure can be the dominant cost factors. Specifically, one can conclude that lactic acid
production should benefit the most from integrated recovery by extraction. However, in this
manuscript no comparison has been made with other configurations such as batch
fermentations with a coupled recovery. According to the current analysis, recovery of
butanol should be performed using adsorption, although it should be noted that extraction
and pervaporation have TMC values relatively close to those of adsorption. The scale of the
process can also have a significant impact on relative costs of each of the product recovery
units. Integrated product recovery can be an important tool when a lower TMC need to be
reached for phenol and should be performed using extraction. Phenol yields on glucose
decrease with increasing phenol concentrations, resulting in an additional beneficial effect
of integrated product recovery.
6.7. Recommendations for model refinement
The costs calculations of the DSP have an accuracy dictated by the applied method and do
not take auxiliary phase regeneration in specific detail into account.
A method to estimate CAPEX costs for DSP regeneration operations is to link the overall
heat duty of the regeneration step to CAPEX requirements [70]. By adding the specific
CAPEX for regeneration operation to the already calculated CAPEX, as describe in this
paper, the regeneration of the auxiliaries, e.g. the solvent phase or adsorption material, can
be taken into account.
Regeneration can be in the form of back-extraction, temperature based regeneration or
distillation. This refinement will further require the calculation of specific heat duties
during regeneration and so product and process dependant properties need to be taken into
account for each separation technique. E.g. butanol recovery from an organic solvent can be
done by distillation and one would operate the distillation column making use of the
existence of an azeotropic point in butanol/water mixtures. Furthermore the OPEX of the
DSP can now be recalculated as more detailed energy requirements are available.
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Appendix A. Nomenclature
volume-specific area
capital expenditure
diffusion coefficient
equipment or material costs
mass-specific enthalpy change
product flow
initial flux
mass transfer coefficient
mass transfer rate
length of hollow fiber
operational expenditure
partition coefficient
adsorbed concentration
mass-specific heat
volumetric-specific productivity
productivity fermentation
particle size
mole fraction
mass-specific work
inhibition constant
axial distance
energy efficiency
solvent phase hold-up
volume flow rate
permeate phase
feed phase
water phase
internal diameter of hollow fiber
theoretical maximum
mixer settler
organic phase
product on substrate
Sauter mean drop-size
Appendix B. Cost-relations for recoverya and fermentation
Fermentation vessel CAPEX  (a  Vferm
 c) . 4.93/12
a = 10035 [euro/m3]
b = 0.608
c = 54840 [euro]
L–L extraction CAPEX mix_set  (a Vmix
_ set ) 
a = 161194 [euro/m3]
b = 0.448
Adsorption CAPEX adsorp =M zeolite  ECzeolite  (a V b  c) 
a = 7650 [euro/m3]
b = 0.347 [due to packing relatively size insensitive]
c =−11525.6 [euro]
Pervaporation and pertraction CAPEX pervap =A mem  ECmem  N units  ECunit 
Capex functions for regeneration of auxiliary phase is assumed in the costs calculations to have the same process time-constants as the
initial dsp operation; to take regeneration into account the calculated dsp capex is doubled if a regeneration step is involved.
Lang-factor breakdown [Ref 75]
1 Major equipment, total purchase cost (PCE)
equipment erection
Buildings, process
site development
ancillary buildings
2. Total physical plant costs (PPC)
PPC = PCE x (1 + f1 +…+ f9)
design and engineering
contractor's fee
Fixed capital = PPC x (1 + f10 + f11 +f12)
Total fixed capital is:
PCE x 3.4 x 1.45 = PCE x 4.93
Appendix C. Cost table
Adsorbent material
Membrane area
Membrane housing
Calcium hydroxide
Hydrochloric acid
Costs euro /unit
Chapter 7: Outlook to bio-based butanol recovery
7.1. Renewable bio-based butanol
1-Butanol has usage as not only a fuel, but also as organic chemical. The world chemicals
consumption is at least 10 times smaller than the world fuel consumption. The strong focus
in society on renewable production of fuels and of bulk chemicals coincides for butanol and
explains the high interest for this compound, as not just an alternative biofuel.
7.2. Butanol from carbohydrates or hydrocarbon
When looking to the future it can be worthwhile to note the past. In the past century, for
commercial production of butanol a competition took place between microbial production
using carbohydrate feedstocks and the newly arising hydrocarbon based petrochemical
industry. This competition was impressively won by the petrochemical industry, becoming
completely market dominant. Multiple market reasons determined the outcome, but the
main factor was the abundance of low priced feedstock, oil. This oil dominance led to the
further development and optimization of oil refineries. When developing new (or old) biobased butanol production processes, one should realize the shear amount of research,
infrastructure and investment made into petrochemical processes. Direct competition will
be hard to almost impossible, as long as the boundary conditions do not change. The most
dramatic boundary condition is the availability of fossil fuels, which will inevitably run out,
but as long as fossil fuels are available their usage will continue.
Major differences exist between petrochemical and microbial based production systems, the
more dramatic being the living nature of microbes. This difference translates into a phase
difference to exist in the production processes, where the microbial systems have a need for
aqueous production vessels (fermentors), and where most conversions in the petrochemical
industry occur in an organic environment, with water being mostly used for heating and
cooling in the form of steam. The aqueous fermentation systems will be predominantly
water with relatively low amounts of organic product. In the case of butanol this can be up
to 30 g/L of butanol product [1]. For less inhibiting biochemicals this can excess 100 g/L.
Crude oil can contain a small water mass fraction, but the organic product content in
petrochemical streams far exceeds fermentation stream product content and can approach
100 %. Due to the volatile nature of most important petrochemical products and the
possibility to separate them based on difference in boiling point, distillation has grown into
the downstream processing method of choice. Distillation requires energy input to enable
the phase change of liquid phase organic compounds to gaseous phase purified products
and is therefore energy intensive. Heat exchange equipment and heat integration methods
allow recovery of substantial amounts of heat at the cost of capital investment. The frequent
use and state of development of distillation as recovery method means that the shear
amount of investments made in distillation and the already existing equipment makes
distillation a logical choice for the recovery of butanol. However, one should realize that
the original design for distillation (excluding liquor production, as it is a added value
consumer commodity) is based primarily on organic feed phase handling. The short term
solution for most butanol recovery problems will be distillation, but the scientific
development of alternative future separation methods is required, as, most definitely, the
basis of the production systems have changed.
7.3. Microbial production of butanol: feedstock
Microbial production of butanol requires feedstock and microbes to convert the feedstock
into desired product, making the matching between feedstock and microbe important. Feed
sugars or so called first generation feedstocks are seen as directly competing with human
food market. Direct usage of these feedstocks is unethical as long as food shortage or high
food prices exist and non-food lignocellulosic material is thus the future feedstock of
choice. As can be seen from the amount of focus on feedstock utilization, feedstocks have
different forms and composition [2,3], and various methods of converting the feedstocks
into substrates for later microbial or chemical conversion [4]. If possible, one would want
to introduce the concentrated raw feedstock without purification into the fermentation
process. The fermentation stability and thus the duration of the fermentation can be
significantly impacted by the presence of inhibiting compounds [5,6]. For most
fermentations a coupling exist between 1) the feedstock and thus feedstock composition 2)
microbial production yield and 3) microbial productivity. The downstream processing can
also be significantly impacted by components present in the feed material, but this has not
been systematically studied yet. Due to these factors, the economic value of butanol
production is a complex function of feedstock properties and therefore local conditions and
geographic considerations determine when and where butanol can be produced
7.4. Microbial 1-butanol production: optimization
Butanol can be produced by bacteria [7] and even yeasts [8]. Besides 1-butanol, 2-butanol
or iso-butanol can be produced. Classic fermentation is by clostridium based AcetoneButanol-Ethanol (ABE) fermentation. Metabolic engineering has focussed on obtaining a
production strain with enhanced solvent tolerance to allow higher titers [9] or enhanced
product yield, e.g. maximizing the amount of butanol produced [1,10]. Production of the
cell material itself limits the overall butanol yield on substrate. Operation close to
maintenance conditions can limit growth and maximize the product yield, as long as
butanol formation leads to net ATP production. By using a cell free synthesis approach,
immobilized enzymes needed for desired conversions are used and thus further biomass
production is avoided; of course such a system can no longer be seen as a fermentative
process [11], but it may be less sensitive to butanol toxicity.
Butanol production (by clostridia) is often associated with butyric acid production,
preceding butanol formation. Producing butanol almost exclusively, avoiding an acidic
production phase altogether, has been achieved using E. coli [1]. However, reaching
unlimited product titers is an unlikely feat for any micro-organism [6] and most probably
for butanol fermentations the actual total solvent concentration will always be below 50
g/L. From a downstream processing perspective it is always beneficial to enhance the
butanol concentration, especially for thermal based recovery methods like distillation. The
largest benefit in reducing downstream processing costs is gained when going from low
concentrations to slightly higher concentrations. Later rates of returns of increases in
concentrations are lower.
One way to limit the build-up of inhibiting compounds in the fermentation broth is to
convert butanol in-situ to a new product. This is feasible when simultaneously performing
an enzymatic reaction in a two phase system between e.g. short chain acid and butanol
[Patent van den Berg, Straathof, v.d. wielen 2011, patent application] or by combining two
1-butanol molecules to form dibutyl ether, which is an unlikely reaction in aqueous
environment. If 1-butanol is the desired product, reaction-based removal is not an option.
7.5. Microbial 1-butanol production: integrated product recovery
1-Butanol is a fermentation inhibiting product, the removal of which will be beneficial for
the fermentation. To that end (continuous) product removal techniques enhance process
productivity, since the fermentation can be maintained [12]. If a batch would have to be
stopped at 20 g/L butanol, and this concentration would be achieved after 10 h, for
example, the downtime of the fermenter might exceed its production time. Only in theory
in-situ product removal could allow indefinite continuous operation of the fermentation. To
date most fermentations are carried out as extended fed-batch like operations. A large array
of similar fermentation units can be scheduled consecutively in such a way that the overall
production process works similar to a continuous plug flow reactor type operation [13]. Our
analysis in chapter 6 shows 1-butanol fermentations, where the aqueous dissolved products
reach modest to low titers, to be a system in which capital investment and operational costs,
for both fermentation as downstream processing, to be similar in overall importance. 1Butanol fermentations thus need to be optimized as an integrated production system in
order to obtain true production cost optimization.
7.6. Recovery of bio-based butanol
Butanol needs to be removed from the fermentation broth and a (direct) phase transition is
needed. However, many phase transitions are associated with a large enthalpy or entropy
The phase transitions shown in chapter 2 show the large (theoretical) list of recovery
options. The odd-one-out is liquid to solid phase transitions as the fusion enthalpy is
substantially lower in value. Basic guidelines and general engineering rules of thumb
unfortunately (the current optimized state of affairs) say solid phase handling is not
desirable, as it makes equipment design far more difficult. Also cooling below ambient
conditions is associated with more operationally expensive (electrical) cooling by
compression, unlike heating which can be done by steam. Not withstanding, freeze
crystallization is, on its merits of enthalpy change, preferred over liquid to gas phase
transition (distillation).
Compared to distillation other recovery techniques such as adsorption, extraction, and
membrane techniques like pervaporation are definitely worthwhile. All listed techniques
can be run in a (semi)-continuous like manner, with secondary product or (inert) feed
component accumulation and fouling of the system being the last hurdle to be taken. E.g.,
thermal based recovery is unable to remove non-volatiles, and extractive process can’t
remove components with a low partitioning in the applied extractive phase. Due to the wide
range of possible feedstock and possible contaminants, robust 1-butanol fermentation
processes will require removal of those specific contaminants when they are (substantially)
inhibiting the fermentation. Also membrane operations can be significantly impacted by
(microbial) fouling of the membranes, leading to significant increases in pressure drop and
an associated increase in operational energy requirement.
Smart choice of the phase manipulation is by constructing the newly formed phase out of
product, for 1-butanol this can be done by phase manipulation of water and butanol
mixtures, creating a binary liquid-liquid system one of which is an organic product phase.
Fermentations that come close to 80 g/L product titers, will be extremely useful when
looking at these liquid-liquid phase separations of water-butanol mixtures. Preferably such
a phase separation is induced by a readily available low cost compound. For fermentations
this can be either carbon dioxide or concentrated feedstock, as feedstock is required
anyway. For longer running operations, especially if internal recycles are applied,
maintaining volumetric flow profiles is difficult when applying low concentration feed
material. Usage of concentrated feed streams thus limits the amount of feed flow into the
production process.
Recovery by pervaporation is possible for butanol, and is mostly constrained by the low
vapour pressure of butanol. Optimization of membrane structure and matrix to optimize
butanol recovery is most often a trade-off between higher product fluxes and lower
selectivity [14]. When this trade-off does not appear directly in reported data, this can most
often be led back to the operation of the recovery system. By operating the system at higher
or lower superficial flow velocities the apparent amount of membrane area can be
manipulated, reporting different selectivities for the same membrane.
Production at
elevated temperature by thermophilic bacteria allows operation at higher temperature,
which positively impacts the saturated vapour pressure of butanol. A large downside of
fermenting at elevated temperatures is additional energy content of the fermentation broth,
although heat can usually be taken from other process steps such as pretreatment,
enzymatic hydrolysis and or sterilization operations, which are mostly carried out above 60
Adsorptive based recovery of butanol from the liquid fermentative phase with temperature
swing regeneration of the sorbent (butanol vaporization) requires the recovery unit to both
experience liquid and gas phase and the unit will experience large density difference
between both phases, besides the possibility of liquid adhesion to the adsorption material,
which can also increase mass transfer problems. Any operation will need to be designed
such a manner that the mechanical operation of having both liquid phase adsorption and gas
phase recovery in the same vessel can be performed without damaging the system. Also this
density difference in phase can pose significant constrains on the sorbent agent.
Sorbent agent thus needs to be durable for both mechanically and temperature aspects.
Adsorptive properties and mass transfer considerations will finally determine the
composition of the basic adsorbent, its overall size and matrix (if e.g. a granulation process
is applied).
For adsorptive based butanol recovery high-silica zeolite has been shown to be highly
selective for butanol compared to water. Also the crystal structure can be chosen in such a
manner that adsorption of feed components, like glucose, is avoided. Decreasing the
hydrophobic nature of the sorbent agents will positively impact temperature based recovery
at the cost of reduced adsorptive properties.
Adsorptive recovery of product from a gas phase auxiliary phase is also possible. This
thesis shows the adsorptive liquid recovery of butanol. By first performing a stripping
operation, butanol can be recovered from a gas phase, avoiding fouling of adsorbent by
non-volatile contaminants and the mechanical stresses of liquid-gas handling in one unit
operation. Recovery by adsorption from the gas phase is limited by the vapour phase
fraction of butanol in the stripping gas (e.g. CO2), which is most often low due to the low
saturated vapour pressure of butanol. This directly leads to a high volumetric stripping gas
requirement. As has also been mentioned for pervaporation, a thermophilic fermentation
will change the gas to liquid equilibria conditions favourably. Gas phase adsorption using
stripping gas approach is more beneficial when applied to intrinsically more suitable
chemicals (defined here as better saturated vapour pressure in most phases), like ethanol.
Currently industrial recovery of bio-butanol is being performed by both Gevo Inc. and
Butamax ™ advanced biofuels LLC. The product of interest is iso-butanol, and is used as
biofuel or a precursor for chemical synthesis.
The applied recovery strategy uses
vaporization under vacuum and a combination of vapour phase manipulation with
absorption and distillation [15]. Conditions for the vacuum, which vaporizes part of the
fermentation liquid, are at a boiling temperature below 30oC, so one can estimate a working
pressure of approx. 500 to 3000 Pa. If the vapour phase contains 10 % butanol, energy duty
calculations at 3000 Pa for vacuum formation, will give 8 MJ/kg. The vacuum step is
followed by distillation also demanding additional energy. Estimates for flash fermentation
of 1-butanol provide an overall energy requirement of 17 MJ/kg butanol, with the vacuum
step demanding 36 % of the energy requirement (6.1 MJ/kg) [16]. The overall energy
requirement for vacuum based recovery like the Butamax proposed process, is close in
range to most of the process options shown in the energy requirement tables 9 and 10 in
chapter 2. Also, like for a standard distillation approach, this process involves only gas and
liquid phase, but can be applied directly on the fermentation liquid, which simplifies the
operation, making the vacuum based recovery a viable industrial recovery method.
7.7. Bio-butanol pricing
To estimate a cost price for bio-based 1-butanol I have increased the detail in the cost price
calculations as shown in Chapter 6, outlined in paragraph 7. If one is able to bring down the
total energy requirement of the recovery to approximately 3.5 MJ/kg of 1-butanol, the dsp
would have a capital expenditure of 0.53 euro/kg and an operational expenditure of approx.
0.05 euro/kg. The fermentation costs are 0.35 and 0.41 euro/kg, respectively. Total 1butanol cost price is 1.34 euro/kg. Finally, it should be noted that the cost price is directly
related to feedstock price and constrained by the butanol concentration during fermentation.
Future developments for feedstock price or 1-butanol fermentation conditions will directly
influence the butanol cost price to a large extend.
7.8. Conclusions
Bio-based butanol production is achievable, as is currently demonstrated by new or
remodelled butanol production facilities in China [17], the U.K. (Gevo Inc.) and the USA
(Butamax advanced biofuels LLC, a joint venture of BP and DuPont). The future will see 1butanol production to be integrated into existing and newly developed productions plants
and dedicated facilities for solely 1-butanol production will not exist.
Butanol production strain development was focused on tolerance and product yield and has
significantly progressed in recent years. In the coming years the focus will be on
fermentation technology, and on usage of impure and mixed 2nd generation feedstocks.
Integrated product removal for 1-butanol is always an optimization problem of fermentative
costs versus downstream processing costs, the phase transition of which will be initially via
gas phase recovery, using distillation, or stripping combined with adsorption.
Butanol is only one of many green based chemicals and has the distinction to have
favourable fuel characteristics and solvent properties. These two product aspects make sure
butanol will be one of the chemicals produced in a future bio-based renewable industry.
The amount of 1-butanol that inevitably will be produced is dependent on the value
associated with butanol.
Reference List
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[2] C. Conde-mejia, A. Jimenez-Gutierrez, El-Halwagi M, A comparison of pretreatment
methods from lignocellulosic materials, Process Safety and Environmental Protection
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[5] P. Dürre, New insights and novel developments in clostridial
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J. Y. Chou, T. Hanai, J. C. Liao, Metabolic engineering of Escherichia coli for 1butanol production, Metabolic Engineering 10 (2008) 305-311.
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Propositions accompanying the thesis:
“Recovery of bio-based butanol”
by Arjan Oudshoorn
1) Research without basic questions leads to a lot of publications.
2) Paraphrasing Benjamin Franklin: “the weakness of democracy is the voter” is more
relevant now than in his day.
3) In today’s specialist society ‘informed consent’ places responsibility in the hands of
the non-specialist.
4) Biofuels should have been named Sustainable Fuels.
5) Humans have great difficulty in coping with probability.
6) Applied Science is a subfield of History.
7) Sustainability in economics leads corporate views to refocus on corporate continuity.
8) Innovation is hampered by conservative business models and risk aversion.
9) A thesis is not a resume.
10) An Engineer should always apply Lefler’s law # 36.
These propositions are regarded as opposable and defendable, and have been approved as
such by the supervisor, Prof. Dr. Ir. L.A.M. van der Wielen.
Stellingen behorende bij het proefschrift:
“Terugwinning van bio-butanol”
door Arjan Oudshoorn
1) Onderzoek zonder fundamentele vragen levert veel publicaties op.
2) Benjamin Franklin geparafraseerd: “de zwakte van een democratie is de kiezer” is
vandaag de dag meer relevant dan in zijn eigen tijd.
3) In de huidige specialistische maatschappij plaatst “toestemming op basis van
informatievoorziening” de verantwoordelijkheid bij de niet-specialist.
4) Biobrandstoffen hadden duurzame brandstoffen moeten heten.
5) Mensen hebben moeite om te gaan met waarschijnlijkheid.
6) Natuurwetenschap is een onderdeel van Geschiedenis.
7) Duurzaamheid in economie zorgt ervoor dat bedrijfsvisies zich weer richten op de
continuïteit van het bedrijf.
8) Innovatie wordt gehinderd door conservatieve bedrijfsmodellen en risicoaversie.
9) Een proefschrift is geen curriculum vitae.
10) Een ingenieur moet altijd Lefler’s Law 36 toepassen.
Deze stellingen worden opponeerbaar en verdedigbaar geacht en zijn zodanig
goedgekeurd door de promotor, Prof. dr. ir. L.A.M. van der Wielen.
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