Large-scale production of oxymethylene dimethyl
ethers
vorgelegt von
Master of Science (M.Sc.)
Franz Kaspar Mantei
an der Fakultät III - Prozesswissenschaften
der Technischen Universität Berlin
zur Erlangung des akademischen Grades
Doktor der Ingenieurwissenschaften
- Dr.-Ing. -
genehmigte Dissertation
Promotionsausschuss:
Vorsitzender: Prof. Dr. Frank Behrendt
Gutachter: Prof. Dr.-Ing. Matthias Kraume
Gutachter: Prof. Dr.-Ing. Erik von Harbou
Tag der wissenschaftlichen Aussprache: 16. Oktober 2023
Berlin 2023
Abstract
Oxymethylene dimethyl ethers (OME) show promising solubility and combustion properties for
applications in various chemical processes and sectors. They enable clean and quasi soot-free
combustion, which can strongly reduce NO
x
emissions. Besides reducing local emissions, OME
can strongly reduce CO
2
emissions by replacing fossil diesel fuel if their production is based
on sustainable methanol. Various process concepts for their production were proposed and
investigated, but most of them prevail signifcant bottlenecks, which prevent their demonstration
and scale-up in the near future. One of the main hurdles is the separation of the by-product
H
2
O which forms in various process steps from H
2
and CO
2
via methanol towards OME
3-5
.
Especially in the OME
3-5
sub-process the separation of H
2
O is challenging considering a large-
scale production plant. Therefore, the novel COMET (clean OME technology) process concept
is introduced and experimentally demonstrated by relying only on state-of-the-art process units.
The COMET process relies solely on methanol and formalin as feedstock and overcomes the
challenging water management using a reactive distillation column.
As the heart of the production process, a suitable catalyst is required to selectively form the
target OME product mixture. Various catalysts have been investigated for OME synthesis focusing
on selectivity and activity. The investigations in this work focus on commercial heterogeneous
catalysts and compare not only the conversions, selectivities and the target product yield but
also the activity, side product formation and thermal stability of the synthesis products. Various
ion exchange resins, zeolites and Nafon catalysts were applied for the OME synthesis in a
batch autoclave at 60
◦C
for the aqueous reaction systems methanol-paraformaldehyde and the
anhydrous reaction system OME
1
-trioxane. Investigations of the synthesis products in a micro
distillation setup showed that all applied catalysts lead to active species in the synthesis product,
negatively impacting its thermal stability. This indicates that a synthesis product handling step
is necessary prior to the downstream purifcation. Based on these investigations, ion exchange
resins are identifed as the most suitable for industrial OME synthesis due to their higher activity
and lower side product formation.
The COMET process and four additional processes for the production of OME
3-5
are simulated
and evaluated, and key performance indicators are defned and compared with alternative
processes from the literature at a scale of 100
kt a�1
OME
3-5
product for the system boundary
starting from H
2
O electrolysis and CO
2
capture. The overall energy efciency for all considered
OME
3-5
production processes is
<
40 % even after heat integration. However, the overall energy
efciency can be signifcantly improved if high temperature heat pumps (HTHP) are used to lift
the temperature level of low temperature excess heat streams. The evaluation shows that by
upgrading excess heat streams using HTHP, a process overall energy efciency of higher than
61 % can be achieved compared to 30 % in a conventional integrated processes. Thereby, the
excess heat stream from H
2
O electrolysis already covers the low temperature heat demand for
CO2 capture via direct air capture, not only for OME but also for various PtX products.
ii
Kurzfassung
Oxymethylendimethylether (OME) weisen vielversprechende Lösungsmittel- und Kraftstofeigen-
schaften für die Verwendung in diversen chemischen Prozessen und Sektoren auf. Sie ermöglichen
eine saubere und nahezu rußfreie Verbrennung, welche die NOx-Emissionen deutlich reduzieren
kann. Neben der Reduzierung lokaler Emissionen können OME durch den Ersatz fossilen Diesel-
kraftstofes die CO
2
-Emissionen bedeutend reduzieren, wenn deren Herstellung auf nachhaltigem
Methanol basiert. Verschiedene Prozesskonzepte für die Herstellung von OME wurden bereits
vorgeschlagen und untersucht, von denen die meisten jedoch noch große Herausforderungen
aufweisen, welche die Demonstration und Maßstabsvergrößerung in naher Zukunft behindern.
Eine der größten Herausforderungen ist die Abtrennung des Nebenproduktes H
2
O, welches in
mehreren Syntheseschritten von H
2
und CO
2
über Methanol und bis hin zu OME
3-5
gebildet wird.
Insbesondere im OME
3-5
Teilprozess ist die Abtrennung des H
2
O schwierig für eine Produktion
im industriellen Maßstab. In dieser Arbeit wird das neue COMET (clean OME technology)
Prozesskonzept vorgestellt, welches experimentell demonstriert wurden und lediglich Prozesskom-
ponenten nach dem aktuellen Stand der Technik beinhaltet. Der COMET Prozess basiert auf
den Edukten Methanol und Formalin und überwindet das herausfordernde Wassermanagement
mit einer Reaktivdestillationskolonne.
Als zentraler Bestandteil des Herstellungsverfahrens wird ein passender Katalysator benötigt,
welcher das OME-Zielprodukt selektiv herstellt. Es wurden bereits zahlreiche Katalysatoren
für die OME-Synthese untersucht, mit dem Fokus auf Selektivität und Aktivität. Die Un-
tersuchungen in dieser Arbeit konzentrieren sich auf kommerzielle, heterogene Katalysatoren.
Neben der Umwandlung, Selektivität und der Zielproduktausbeute wurden auch die Aktivität,
Nebenproduktbildung und thermische Stabilität des Syntheseproduktes untersucht. Mehrere
Ionenaustauschharze, Zeolithe und Nafon wurden als Katalysatoren für die OME-Synthese in
einem Batch-Autoklaven bei 60
◦C
für das wässriges Reaktionssystem Methanol-Paraformaldehyd
und für das wasserfreie Reaktionssystem OME
1
-Trioxan eingesetzt. Untersuchungen der Synthe-
seprodukte in einer Labordestillationsanlage haben gezeigt, dass alle Katalysatoren zu aktiven
Zentren im Syntheseprodukt führen, welche die thermische Stabilität negativ beeinfussen. Das
deutet darauf hin, dass eine Nachbehandlung des Syntheseproduktes vor der anschließenden
Auftrennung notwendig ist. Die Ergebnisse zeigen, dass Ionenaustauschharze aufgrund ihrer hohen
Aktivität und geringen Nebenproduktbildung am geeignetsten für die industrielle OME-Synthese
sind.
Der COMET Prozess und vier weitere Prozesse zur Herstellung von OME
3-5
wurden mit
einer Kapazität von 100
kt a�1
OME
3-5
basierend auf H
2
von einer Wasserelektrolyse und
abgeschiedenem CO
2
, simuliert. Wichtige Leistungsindikatoren wurden defniert und mit alterna-
tiven Herstellunsgsverfahren aus der Literatur verglichen. Die Gesamtenergieefzienz beträgt
bei allen OME
3-5
Herstellungsverfahren
<
40 %, auch nach der Wärmeintegration. Jedoch kann
die Gesamtenergieefzienz durch den Einsatz von Hochtemperaturwärmepumpen für die Tempe-
raturerhöhung von Niedertemperaturüberschusswärmeströmen deutlich verbessert werden. Die
iv
Kurzfassung
Ergebnisse zeigen, dass durch den Einsatz der Hochtemperaturwärmepumpen die Gesamtener-
gieefzienz von etwa 30 % auf über 61 % angehoben werden kann. Dabei stellt die Abwärme
der Wasserelektrolyse bereits genügend Wärme für den Niedertemperaturwärmebedarf der CO
2
Abtrennung direkt aus der Luft bereit. Dies gilt nicht nur für OME, sondern für alle betrachteten
PtX Produkte.
Die experimentelle Validierung, Simulationsarbeit und methodische Auswertung dieser Arbeit
ebnen den Weg für weiterführende Planungen und Umsetzungen industrieller OME Herstellungs-
verfahren.
v
Contents
Abstract ............................................. ii
Kurzfassung ........................................... iv
Notation ............................................. ix
List of Figures .......................................... xiii
List of Tables ........................................... xviii
1 Introduction ......................................... 1
1.1 Context and Background ............................... 1
1.2 Scope and Objectives of this thesis .......................... 4
2 State of the Art ....................................... 8
2.1 OME Properties and Applications .......................... 8
2.2 OME synthesis ..................................... 9
2.3 Catalyst systems for the synthesis of OME . . . . . . . . . . . . . . . . . . . . . . 11
2.4 OME3-5 production processes ............................. 13
2.5 H2O separation from the production of OME . . . . . . . . . . . . . . . . . . . . 14
2.5.1 Extraction.................................... 15
2.5.2 Adsorption ................................... 16
2.5.3 Membrane.................................... 17
3 Methods ........................................... 20
3.1 Experimental investigation of the OME synthesis . . . . . . . . . . . . . . . . . . 20
3.1.1 Chemicalsandmaterials............................ 20
3.1.2 Analytics .................................... 20
3.1.3 Apparatus.................................... 21
3.1.4 Feed preparation ................................ 21
3.1.5 OME synthesis ................................. 22
3.1.6 Catalyst performance evaluation ....................... 23
3.1.7 Distillation of the OME synthesis products . . . . . . . . . . . . . . . . . 24
3.2 Experimental demonstration of the main COMET process units . . . . . . . . . . 25
3.2.1 Chemicalsandmaterials............................ 25
3.2.2 Analytics .................................... 25
3.2.3 Concentrated FA(aq.) feed preparation . . . . . . . . . . . . . . . . . . . 25
3.2.4 OME synthesis ................................. 26
3.2.5 OME synthesis product neutralization . . . . . . . . . . . . . . . . . . . . 26
3.2.6 Thermal separation in CO-1 . . . . . . . . . . . . . . . . . . . . . . . . . . 27
3.2.7 Reactive distillation in CO-2 ......................... 27
vi
Contents
3.2.8 Thermal separation in CO-3 . . . . . . . . . . . . . . . . . . . . . . . . . . 27
3.3 Process simulation and evaluation .......................... 29
3.3.1 General assumptions and system boundaries . . . . . . . . . . . . . . . . . 29
3.3.2 Process modelling and simulation . . . . . . . . . . . . . . . . . . . . . . . 29
3.3.3 Process evaluation and comparison criteria . . . . . . . . . . . . . . . . . 31
3.4
Process evaluation for an improved energy efciency of Power-to-X processes using
heat pumps ....................................... 32
4 OME process description .................................. 35
4.1 OME3-5 production processes ............................. 35
4.1.1 OME1 and TRI (anhydrous synthesis) .................... 36
4.1.2 DME and TRI (anhydrous synthesis) . . . . . . . . . . . . . . . . . . . . . 38
4.1.3 DME and monomeric FA (anhydrous synthesis) . . . . . . . . . . . . . . . 39
4.1.4 OME1 and monomeric FA (anhydrous synthesis) . . . . . . . . . . . . . . 40
4.1.5 MeOH and FA(aq.) (aqueous synthesis) ................... 41
4.1.6 MeOH and monomeric FA (aqueous synthesis) . . . . . . . . . . . . . . . 42
4.1.7 OME1 and FA(aq.) or pFA (aqueous synthesis) . . . . . . . . . . . . . . . 43
4.1.8 COMET (aqueous synthesis) ......................... 44
4.1.9 Further process concepts for the production of OME3-5 .......... 47
4.2 Upstream processes for the OME production based on H2 and CO2 ........ 48
4.2.1 MeOH production ............................... 48
4.2.2 FA(aq.) production and concentration .................... 49
4.2.3 monomericFAproduction........................... 50
4.2.4 OME1 production ............................... 51
4.2.5 Combustion................................... 51
5 Results and Discussion ................................... 54
5.1 Experimental investigation of the OME synthesis . . . . . . . . . . . . . . . . . . 54
5.1.1 Reaction progress and equilibrium composition . . . . . . . . . . . . . . . 54
5.1.2 Catalyst activity ................................ 57
5.1.3 Side and by-product formation ........................ 58
5.1.4 Thermal stability of the synthesis products . . . . . . . . . . . . . . . . . 63
5.2 Experimental demonstration of the main COMET process units . . . . . . . . . . 66
5.2.1 OME synthesis ................................. 66
5.2.2 Synthesis product neutralization ....................... 67
5.2.3 Synthesis product separation in CO-1 . . . . . . . . . . . . . . . . . . . . 68
5.2.4 Reactive distillation in CO-2 ......................... 69
5.2.5 Product separation in CO-3 .......................... 70
5.3 Process simulation ................................... 72
5.3.1 Processdescription............................... 72
5.3.2 Mass balance .................................. 77
5.3.3 Energydemand................................. 79
vii
Table of contents
5.3.4 Process efciencies ............................... 81
5.3.5 Comparison of alternative OME3-5 production processes . . . . . . . . . . 82
5.4
Process evaluation for an improved energy efciency of Power-to-X processes using
heat pumps ....................................... 85
5.4.1 OME3-5 production from H2 and CO2 (system boundary I) . . . . . . . . 85
5.4.2 Including H2 production (system boundary II) . . . . . . . . . . . . . . . 86
5.4.3 Including CO2 preparation (system boundary III) . . . . . . . . . . . . . 88
5.4.4 Including HTHP systems (system boundary IV) . . . . . . . . . . . . . . 90
5.4.5 Potential applications for the recovered excess heat (system boundary V) 91
5.4.6
Potential analysis: Upgrading the excess heat of the H
2
O electrolysis using
HTHP to supply the low-temperature heat demand of DAC systems for
various PtX products ............................. 92
6 Conclusion and outlook ................................... 95
References ............................................ 99
Acknowledgement ........................................ 117
Appendix ............................................. A
A.1 Experimental investigation of the OME synthesis . . . . . . . . . . . . . . . . . . A
A.1.1 Reaction progress and equilibrium composition . . . . . . . . . . . . . . . A
A.1.2 Thermal stability of the synthesis products . . . . . . . . . . . . . . . . . R
A.2 Experimental demonstration of the main COMET process units . . . . . . . . . . T
A.2.1 OME synthesis ................................. T
A.2.2 Synthesis product separation in CO-1 . . . . . . . . . . . . . . . . . . . . U
A.2.3 Reactive distillation in CO-2 ......................... U
A.2.4 Product separation in CO-3 .......................... U
A.3 Process modelling and simulation . . . . . . . . . . . . . . . . . . . . . . . . . . . W
A.3.1 Pure component properties .......................... W
A.3.2 Thermodynamic model for mixtures ..................... W
A.3.3 Validation.................................... X
A.3.4
Mass balance and operation conditions of the process units of the COMET
process starting from H2 and CO2 ...................... AA
A.3.5 Comparison to alternative OME3-5 production processes . . . . . . . . . . AK
viii
Notation
Abbreviations
A15 Amberlyst® 15
A36 Amberlyst® 36
A46 Amberlyst® 46
AEL alkaline electrolysis
BV back pressure valve
CO carbon monoxide
CO2 carbon dioxide
CO distillation column
COMET clean OME technology
COP coefcient of performance
CPS carbon from point sources
DAC direct air capture
DBU german federal environmental foundation
DH district heating
DME dimethyl ether
Dowex Dowex® 50WX2
EA ethyl acetate
EGR exhaust gas recirculation
EPDM ethylene propylene diene monomer
EtOH ethanol
F flter
FA formaldehyde
FA(aq.) aqueous FA solution, formalin
FFKM perfuoroelastomer
FOAC formic acid
FT Fischer-Tropsch
GC-FID gas chromatograph equipped with a fame ionization detector
GC-TCD gas chromatograph equipped with a thermal conductivity detector
GHG greenhouse gases
GHGE greenhouse gas emissions
H heat exchanger
H2 hydrogen
H2O water
HF poly(oxymethylene) hemiformals
HT high temperature
HTHP high temperature heat pump
ix
Notation
HVO hydrogenated vegetable oil
IEA international energy agency
IER ion exchange resin
KPIs key performance indicators
LCA life cycle assessment
LHV lower heating value
LPS low pressure steam
MED multi-efect distillation
MEFO methyl formate
MeOH Methanol
MG poly(oxymethylene) glycols
MPS medium pressure steam
MSF multi-stage fash desalination
MtG methanol to gasoline
N2 nitrogen
Nafon Nafon™ NR40
NAMOSYN sustainable mobility through synthetic fuels
NOx nitrous oxides
O2 oxygen
OME oxymethylene dimethyl ethers
OME1 methylal
p product fask
PEM polymer electrolyte membrane
pFA paraformaldehyde
PI pressure indicator
PTFE polytetrafuoroethylene
PtL power-to-liquid
PtX power-to-X
R reactor
RO reverse osmosis
S separator, catalyst chamber, sample
SMR steam methane reforming
SOEC solid oxide electrolyser cell
THF tetrahydrofuran
TI, TIC temperature indicator
TRI trioxane
TRL technology readiness level
TSA temperature-vacuum swing adsorption
V valve
VHTHP very high temperature heat pumps
VLE vapor-liquid equilibria
VLLE vapor-liquid-liquid equilibria
x
Notation
WtW well-to-wheel
Symbols (Latin)
Symbol Designation Unit
A area 2
m
C number of carbon atoms -
d diameter m
H energy content based on the LHV kW
l
LHV
length
lower heating value
m
�1
kW h kg
m
m, ˙mi
mass
mass fow rate
g
h�1
kg
p
˙
Qk
pressure
heat fow
bar
kW
S selectivity
mol mol�1
t time min
T temperature ◦C
V volume ml
wi mass fraction wt%
Wl
xi
eletric power
mole fraction
kW
mol mol�1
X
Y
conversion
yield
�1
g g
�1
g g
Symbols (Greek)
Symbol Designation Unit
η efciency %
Indices
xi
Notation
Symbol Designation
C carbon
cat catalyst
el electric
eq equilibrium
i reactant, product, component, inner
th thermal
xii
List of Figures
1.1 Impact of the anthropogenic climate change at 1.5 and 2 ◦C of warming by [5]. . 1
1.2
The PtX value chain from sustainable feedstocks via efcient conversion processes
towards advanced PtX products. ©Fraunhofer ISE [10] . . . . . . . . . . . . . . . 2
1.3 Schematical overview of the chronology of the publications this thesis is based on. 5
2.1
H
2
O separation from a sustainable production of OME
3-5
based on H
2
and CO
2
[4].
15
2.2
H
2
O separation from the OME synthesis product using toluene as extractant
(conditions: (OME synthesis product)/toluene = 0
.
667
g g�1
, 25
◦C
, batch) by
Li et al. [4, 112]. .................................... 16
2.3
H
2
O separation from the OME synthesis product via adsorption using zeolite 3A
(conditions: (OME mixture)/(zeolite 3A) = 2
.
0
g g�1
, 25
◦C
, batch) by Schmitz
et al. [4, 18]. ...................................... 17
2.4
H
2
O separation from the OME synthesis product using the polymeric membrane
PERVAP 4100 (conditions: 80
◦C
,2
mbar
permeate, 80
l h�1
) by Schmitz et al.
[4, 92]. .......................................... 18
3.1
Autoclave (a) and distillation setup (b) for the investigation of the OME synthesis
and thermal stability of the OME synthesis product [3]. BV, back pressure valve;
F, flter; H, heat exchanger; P, product fask; PI, pressure indicator; S, catalyst
chamber; TI and TIC, temperature indicators; V, valve. . . . . . . . . . . . . . . 22
3.2
Simplifed process fow diagram for the concentration of an FA(aq.) to 85
�
88
wt
%
FA using a cascade of two evaporators operated at 200
�
600
mbar
and 100
�
150
◦C
heating fuid [4]. E, evaporator; P, pump. ...................... 25
3.3
Simplifed process fow diagram for the OME synthesis of OME
1
and concentrated
FA(aq.) solution over A46 for a capacity of 1
�
5
L h�1
at about 90
◦C
[4]. H,
heat exchanger; P, pump; R, reactor; T, tank; V, valve. . . . . . . . . . . . . . . 26
3.4
Simplifed process fow diagram of the DN 50 glass distillation setup for a feed
rate of 1
�
5
L h�1
[4]. CO, distillation column; H, heat exchanger; P, pump; VP,
vacuum pump. ..................................... 27
3.5
System boundaries set on the simulation level of the OME
3-5
production processes
[1]............................................. 29
3.6
System boundary of a sustainable OME
3-5
production process from H
2
and CO
2
for the evaluation of the overall energy efciency. In scenario I, the production
of OME
3-5
is considered from H
2
and CO
2
via the intermediate production of
MeOH and FA, including heat integration. Scenarios II and III additionally cover
the preparation and provision of H
2
via H
2
O electrolysis and CO
2
via DAC. In
scenario IV, HTHP are included to lift low-temperature excess heat streams to
usable temperature levels. Scenario V additionally considers seawater desalination
and scenario VI the application of excess heat streams for other applications [2]. 33
xiii
List of Figures
4.1
OME
3-5
production process for various feedstocks, following aqueous and anhydrous
reaction systems [4]. CO, distillation column; R, reactor; S, H2O separator. . . . 36
4.2
OME synthesis from OME
1
and TRI over A46 (conditions: OME
1
/TRI = 2
.
85
g g�1
,
A46/(OME
1
+TRI) = 0
.
8
wt
%, 65
◦C
, batch) by Burger [4, 14]. (a) shows the
reaction progress and (b) shows the equilibrium composition. The values describe
the mass fractions of the synthesis products. . . . . . . . . . . . . . . . . . . . . . 38
4.3
OME synthesis from DME and TRI over A36 (conditions: DME/TRI = 1
.
80
g g�1
,
A36/(DME+TRI) = 4
.
4
wt
%, 90
◦C
, batch) by Breitkreuz [4, 132]. (a) shows
the reaction progress and (b) shows the composition after 76
h
with the highest
concentration of OME
3-5
. The values describe the mass fractions of the synthesis
products. ........................................ 39
4.4
OME synthesis from OME
1
and monomeric FA over OMe
3+
BF
4-
in EMIM
+
BF
4-
(ionic liquid) (conditions: OME
1
/FA = 1
.
58
g g�1
, OMe
3+
BF
4-
/OME
1
= 2
�
3
mol
%, 45
◦C
, continuous addition of gaseous FA) by Peter et al. [4, 141]. The
values describe the mass fractions of the synthesis product. . . . . . . . . . . . . 40
4.5
OME synthesis from MeOH and pFA over A46 (conditions: FA/MeOH = 0
.
89
g g�1
,
A46/(MeOH+pFA) = 1
.
9
wt
%, 60
◦C
, batch) by Schmitz et al. [4, 80]. (a) shows
the reaction progress and (b) shows the equilibrium composition. The values
describe the mass fractions of the synthesis products. . . . . . . . . . . . . . . . 42
4.6
OME synthesis from OME
1
and pFA over NKC-9 (conditions: OME
1
/pFA
= 4
.
31
g g�1
, NKC-9/(OME
1
+pFA) = 5
wt
%, 3h, 90
◦C
, batch) by Liu
et al. [4, 140]. The values describe the mass fractions of the synthesis product. . 43
4.7
COMET process concept for the production of OME
3-5
from MeOH and FA(aq.)
feedstocks [4]. The light grey arrows and process units were added in this work to
the FA concentration sub-process to improve the recycle of FA. CO, distillation
column; E, evaporator; R, reactor. .......................... 45
4.8
H
2
O separation from the COMET process via reactive distillation [4]. The left
side shows the reactive distillation column with the main components of the feed
and product streams. The illustration on the right side shows the interaction on a
catalytic tray and was adopted from Schmitz et al. [91]. . . . . . . . . . . . . . . 46
4.9
COMET process concept for the production of OME
3-5
from H
2
and CO
2
feedstock
with the intermediate production of MeOH and FA [4]. . . . . . . . . . . . . . . . 47
4.10
Simplifed process fow diagram for the production of MeOH from H
2
and CO
2
based on [4, 107]. .................................... 49
4.11
Simplifed process fow diagram for the production (A) and concentration (B) of
FA(aq.) from MeOH and air based on [4, 109, 134]. . . . . . . . . . . . . . . . . . 50
4.12
Simplifed process fow diagram for the production of monomeric FA from MeOH
based on [4, 110, 128]. ................................. 51
4.13
Simplifed process fow diagram for the production of OME
1
from MeOH and
FA(aq.) based on [4, 129]. ............................... 51
4.14
Simplifed process fow diagram for the combustion of the purge streams based on
[4]............................................. 52
xiv
List of Figures
5.1
OME synthesis from MeOH-pFA over A36 (conditions: pFA/MeOH = 1
.
53
g g�1
,
A36/(MeOH + pFA) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch) [3]. (a) illustrates the
reaction progress and (b) the equilibrium composition after 24 h.......... 54
5.2
OME synthesis from OME
1
-TRI over A36 (conditions: OME
1
/TRI = 2
.
00
g g�1
,
A36/(OME
1
+ TRI) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch) [3]. (a) illustrates the
reaction progress and (b) the equilibrium composition after 24 h.......... 55
5.3
OME
3-5
yield over the synthesis progress for various catalysts (conditions: pFA/MeOH
= 1
.
5
g g�1
, OME
1
/TRI = 2
.
0
g g�1
, catalyst/reactants = 1
.
0
wt
%, 60
◦C
,
8
bar
, batch) [3]. (a) illustrates the results for the MeOH-pFA feed mixture and
(b) illustrates the results for the OME1-TRI feed mixture. ............. 55
5.4
Conversion of the reactants and selectivity towards OME
3-5
for the OME synthesis
from MeOH-pFA and OME
1
-TRI over various catalysts (conditions: pFA/MeOH
= 1
.
5
g g�1
, OME
1
/TRI = 2
.
0
g g�1
, catalyst/reactants = 1
.
0
wt
%, 60
◦C
,8
bar
,
24
h
, batch) [3]. (a) illustrates the results for the MeOH-pFA feed mixture and
(b) the results for the OME1-TRIfeedmixture.................... 56
5.5 Termination time (a) and yield Y ∗ (b) of the OME synthesis from MeOH-
OME3�5
pFA and OME
1
-TRI for various catalysts (conditions: pFA/MeOH = 1
.
5
g g�1
,
OME1/TRI = 2.0 g g�1, catalyst/reactants = 1.0 wt%, 60 ◦C,8 bar, batch) [3]. 57
5.6
MEFO side product formation over the synthesis progress from MeOH-pFA
for various catalysts (conditions: pFA/MeOH = 1
.
5
g g�1
, catalyst/reactants
= 1
.
0
wt
%, 60
◦C
,8
bar
, batch) [3]. The dashed lines show the termination time
oftherespectivecatalysts................................ 58
5.7
MEFO side product formation over the synthesis progress from OME
1
-TRI for
various catalysts (conditions: OME
1
/TRI = 2
.
0
g g�1
, catalyst/reactants =
1
.
0
wt
%, 60
◦C
,8
bar
, batch) [3]. The dashed lines show the termination time of
the respective catalysts. ................................ 60
5.8
TRI side product formation over the synthesis progress from MeOH-pFA for various
catalysts (conditions: pFA/MeOH = 1
.
5
g g�1
, catalyst/reactants = 1
.
0
wt
%,
60
◦C
,8
bar
, batch) [3]. The dashed lines show the termination time of the
respective catalysts. .................................. 61
5.9
FA (a), MeOH (b) and H
2
O (c) side product formation over the synthesis progress
from OME
1
-TRI for various catalysts (conditions: OME
1
/TRI = 2
.
0
g g�1
,
catalyst/reactants = 1.0 wt%, 60 ◦C,8 bar, batch) [3]. ............... 62
5.10 Synthesis product composition and added up distillation product composition of
the OME synthesis from MeOH-pFA (a) and OME
1
-TRI (b) for various catalysts
(conditions synthesis: pFA/MeOH = 1
.
5
g g�1
, OME
1
/TRI = 2
.
0
g g�1
, cata-
lyst/reactants = 1
.
0
wt
%, 60
◦C
,8
bar
, batch; conditions distillation: 30
�
50
g
synthesis product, TOil = 60 � 100 ◦C stepwise, 5 h, batch) [3]. .......... 64
xv
List of Figures
5.11
OME synthesis from OME
1
and concentrated FA(aq.) over A46 (conditions:
concentrated FA(aq.) with 85
�
89
wt
% FA, (concentrated FA(aq.))/OME
1
= 0
.
6
g g�1
, A46/(OME
1
+concentrated FA(aq.)) = 0
.
34
gh g�1
, approx.
3
L h�1
, 90
◦C
, 10
bar
, fxed bed reactor) [4]. F represents the feed composition.
P-Sim, P1-5-Exp represent the product composition of the simulated equilibrium,
two experimental preliminary products P1-2 from the starting phase and the three
product barrels P3-5, respectively. .......................... 66
5.12
CO-1, OME synthesis product separation (conditions: 2
L h�1
, refux/distillate
= 0
.
5
�
2
s s�1
, distillate/feed = 81
wt
%, Montz 750 structured packing,
85
�
175
◦C
, ambient pressure) [4]. The values describe the mass fractions of the
feed mixture, here P5-Exp as presented in Figure 5.11, the distillate product and
bottom product. .................................... 69
5.13
CO-2, reactive distillation of the distillate product of CO-1 over A46 (conditions:
A46/(feed stream) = 0
.
35
gh g�1
,1
L h�1
, distillate/feed = 63
wt
%, Montz 750
structured packing, 45
�
104
◦C
, ambient pressure) [4]. The values describe the
mass fractions of the feed mixture, the distillate product and bottom product. . . 70
5.14
CO-3, product separation (conditions: 5
.
5
L h�1
, distillate/feed = 82
wt
%, Montz
750 structured packing, 100
�
210
◦C
, 200
mbar
) [4]. The values describe the
mass fractions of the feed mixture, here the CO-1 bottom product, the distillate
product and bottom product. ............................. 71
5.15
Simplifed process fow diagram of P1 for the production of OME
3-5
from H
2
and
CO2 via the intermediate production of MeOH and FA(aq.) [1]. . . . . . . . . . . 72
5.16
Simplifed process fow diagram of P2 for the production of OME
3-5
from H
2
and
CO2 via the intermediate production of MeOH and monomeric FA [1]. . . . . . . 73
5.17
Simplifed process fow diagram of P3 for the production of OME
3-5
from H
2
and
CO2 via the intermediate production of MeOH, FA(aq.) and OME1 [1]. ..... 74
5.18
Simplifed process fow diagram of P4 for the production of OME
3-5
from H
2
and
CO2 via the intermediate production of MeOH, monomeric FA and OME1 [1]. . 75
5.19
Process energy efciency progress for producing OME
3-5
from H
2
and CO
2
includ-
ing H
2
O electrolysis, seawater desalination, and CO
2
capture directly from air.
The process is heat-integrated, and low-temperature product streams are lifted to
100
◦C
using HTHP and supply the heat demand for DAC, seawater desalination,
and other applications [2]. ............................... 87
5.20
Utilization of the low temperature product heat stream from the OME
3-5
produc-
tion process and H2O electrolysis upgraded using HTHP [2]. . . . . . . . . . . . . 92
A.1
OME synthesis from MeOH-pFA over (a) A15, (b) A36, (c) A46, (d) Dowex,
(e) H-BEA25, (f) H-MFI90, (g) Nafon (conditions: pFA/MeOH = 1
.
50
g g�1
,
catalyst/(MeOH+pFA) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch). The termination time
represents the time after which 90 % of the OME
5
concentration after 24
h
are
obtained. ........................................ B
xvi
List of fgures
A.2
OME synthesis from OME
1
-TRI over (a) A15, (b) A36, (c) A46, (d) Dowex,
(e) H-BEA25, (f) H-MFI90, (g) Nafon (conditions: OME
1
/TRI = 2
.
00
g g�1
,
catalyst/(OME
1
+TRI) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch). The termination time
represents the time after which 90 % of the OME
5
concentration after 24
h
are
obtained. ........................................ J
A.3
Bottom product composition and conditions of the distillations of the OME
synthesis products from MeOH-pFA feed mixture. . . . . . . . . . . . . . . . . . S
A.4
Bottom product composition and conditions of the distillations of the OME
synthesis products from OME1-TRI feed mixture. . . . . . . . . . . . . . . . . . . T
A.5 Reft of UNIFAC interaction parameters. . . . . . . . . . . . . . . . . . . . . . . . X
A.6
Average deviation of VLE data from diferent sub-systems as presented in Table
A.23 ........................................... Z
A.7
Simplifed process diagram for the production of MeOH from H
2
and CO
2
. CO,
distillation column; FL, phase separator; HE, heat exchanger; PC, compressor; R,
reactor. ......................................... AA
A.8
Simplifed process diagram for the production of FA(aq.) from MeOH and air.
CO, distillation column; HE, heat exchanger; PC, compressor; R, reactor. . . . . AD
A.9
Simplifed process diagram of the FA concentration based on FA(aq.) solution.
CO, distillation column; E, evaporator; HE, heat exchanger. . . . . . . . . . . . . AG
A.10
Simplifed process diagram for the production of OME
3-5
from MeOH and con-
centrated FA(aq.) solution. CO, distillation column; HE, heat exchanger; R,
reactor. ......................................... AH
A.11
Simplifed process diagram for the combustion of the purge streams. HE, heat
exchanger; PC, compressor; R, reactor. . . . . . . . . . . . . . . . . . . . . . . . . AJ
xvii
List of Tables
2.1 Properties of OME1-6, diesel fuel and HVO [35–40]. . . . . . . . . . . . . . . . . . 8
2.2 Catalysts for the OME synthesis from MeOH-pFA as well as OME1-TRI [1] . . . 14
2.3
Main advantages and main hurdles of the H
2
O separation methods extraction,
adsorption, and membrane [4]. ............................ 18
4.1
Methyl group suppliers and oxymethylene group suppliers for anhydrous and
aqueous OME reaction systems [4]. .......................... 35
4.2 Advantages and main hurdles of various OME3-5 production process concepts [4]. 37
5.1
Stream compositions and conditions of the COMET process presented in Figure
4.7 [4]. .......................................... 76
5.2
Overall mass balance for the production of OME
3-5
following the COMET process
and the processes P1 to P4. The processes were simulated with a capacity of
100 kt a�1 OME3-5 [4].................................. 78
5.3
Overall energy demand for the production of OME
3-5
following the COMET
process and the processes P1 to P4. The processes were simulated with a capacity
of 100 kt a�1 OME3-5 [4]. ............................... 80
5.4
Overall process efciencies for the production of OME
3-5
following the COMET
process and the processes P1 to P4. The processes were simulated with a capacity
of 100 kt a�1 OME3-5 [4]. ............................... 81
5.5
Comparison of various OME
3-5
production processes based on the results of this
work [4], Held et al. [109] and Schemme et al. [119]. ................ 83
5.6
Key performance indicators of the process for the production of 100
kt a�1
OME
3-5
from H2 and CO2 via MeOH and FA(aq.) [1]. .................... 86
5.7
Some commercially available H
2
O electrolysis systems and their energy and H
2
O
demand [2]. ....................................... 88
5.8
Key operation parameters of the H
2
production via H
2
O electrolysis considered in
this work [2]. ...................................... 88
5.9 Some commercially available DAC systems and their specifc energy demands [2]. 89
5.10 Some commercially available HTHP systems and their performance [2]. . . . . . 90
5.11
Waste heat recovery using HTHP systems and their infuence on the overall process
energy efciency, assuming a constant COP of 4
.
5 for temperature lifts from down
to 40 to 100 ◦C [2].................................... 91
5.12
H
2
and CO
2
demand for various PtX products and the recovery of the excess heat
H2
from a PEM or AEL electrolyzer with 10
.
4
kWh kg�1
using HTHP with a COP
of 4
.
5 to supply the heat demand of the DAC system of 1
.
5
kWh kg�1
[1, 2, 119].
93
CO2
xviii
List of Tables
A.1
OME synthesis from MeOH-pFA over A15 (conditions: pFA/MeOH = 1
.
50
g g�1
,
A15/(MeOH+pFA) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch). The concentrations are
presented in mass fractions and the duration in minutes. . . . . . . . . . . . . . . C
A.2
OME synthesis from MeOH-pFA over A36 (conditions: pFA/MeOH = 1
.
53
g g�1
,
A36/(MeOH+pFA) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch). The concentrations are
presented in mass fractions and the duration in minutes. . . . . . . . . . . . . . . D
A.3
OME synthesis from MeOH-pFA over A46 (conditions: pFA/MeOH = 1
.
50
g g�1
,
A46/(MeOH+pFA) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch). The concentrations are
presented in mass fractions and the duration in minutes. . . . . . . . . . . . . . . E
A.4
OME synthesis from MeOH-pFA over Dowex (conditions: pFA/MeOH = 1
.
50
g g�1
,
Dowex/(MeOH+pFA) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch). The concentrations are
presented in mass fractions and the duration in minutes. . . . . . . . . . . . . . . F
A.5
OME synthesis from MeOH-pFA over H-BEA 25 (conditions: pFA/MeOH =
1
.
50
g g�1
, H-BEA 25/(MeOH+pFA) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch). The
concentrations are presented in mass fractions and the duration in minutes. . . . G
A.6
OME synthesis from MeOH-pFA over H-MFI 90 (conditions: pFA/MeOH =
1
.
50
g g�1
, H-MFI 90/(MeOH+pFA) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch). The
concentrations are presented in mass fractions and the duration in minutes. . . . H
A.7
OME synthesis from MeOH-pFA over Nafon (conditions: pFA/MeOH = 1
.
49
g g�1
,
Nafon/(MeOH+pFA) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch). The concentrations are
presented in mass fractions and the duration in minutes. . . . . . . . . . . . . . . I
A.8
OME synthesis from OME
1
-TRI over A15 (conditions: OME
1
/TRI = 1
.
99
g g�1
,
A15/(OME
1
+TRI) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch). The concentrations are
presented in mass fractions and the duration in minutes. . . . . . . . . . . . . . . K
A.9
OME synthesis from OME
1
-TRI over A36 (conditions: OME
1
/TRI = 2
.
00
g g�1
,
A36/(OME
1
+TRI) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch). The concentrations are
presented in mass fractions and the duration in minutes. . . . . . . . . . . . . . . L
A.10
OME synthesis from OME
1
-TRI over A46 (conditions: OME
1
/TRI = 2
.
00
g g�1
,
A46/(OME
1
+TRI) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch). The concentrations are
presented in mass fractions and the duration in minutes. . . . . . . . . . . . . . . M
A.11
OME synthesis from OME
1
-TRI over Dowex (conditions: OME
1
/TRI = 2
.
01
g g�1
,
Dowex/(OME
1
+TRI) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch). The concentrations are
presented in mass fractions and the duration in minutes. . . . . . . . . . . . . . . N
A.12
OME synthesis from OME
1
-TRI over H-BEA 25 (conditions: OME
1
/TRI =
2
.
01
g g�1
, H-BEA 25/(OME
1
+TRI) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch). The
concentrations are presented in mass fractions and the duration in minutes. . . . O
A.13
OME synthesis from OME
1
-TRI over H-MFI 90 (conditions: OME
1
/TRI =
2
.
00
g g�1
, H-MFI 90/(OME
1
+TRI) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch). The
concentrations are presented in mass fractions and the duration in minutes. . . . P
A.14
OME synthesis from OME
1
-TRI over Nafon (conditions: OME
1
/TRI = 1
.
99
g g�1
,
Nafon/(OME
1
+TRI) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch). The concentrations are
presented in mass fractions and the duration in minutes. . . . . . . . . . . . . . . Q
xix
List of Tables
A.15
Bottom product composition and conditions of the distillations of the OME syn-
thesis products from MeOH-pFA feed mixture. The concentrations are presented
in mass fractions. .................................... R
A.16
Bottom product composition and conditions of the distillations of the OME
synthesis products from OME
1
-TRI feed mixture. The concentrations are presented
in mass fractions. .................................... S
A.17
OME synthesis from OME
1
and concentrated FA(aq.) solution over A46 (con-
ditions: concentrated FA(aq.) with 85
�
89
wt
% FA, (concentrated FA(aq.)
solution)/OME
1
= 0
.
6
g g�1
, A46/(OME
1
+concentrated FA(aq.) solution)
= 0
.
34
gh g�1
, approx. 3
L h�1
, 90
◦C
, 10
bar
, fxed bed reactor). The
concentrations are presented in mass fractions. . . . . . . . . . . . . . . . . . . . T
A.18
CO-1, OME synthesis product separation (conditions: 2
L h�1
, refux/distillate
= 0
.
5
�
2
s s�1
, distillate/feed = 81
wt
%, Montz 750 structured packing,
85
�
175
◦C
, ambient pressure). The concentrations are presented in mass fractions.
U
A.19
CO-2, Reactive distillation of the distillate product of CO-1 over A46 (conditions:
A46/(feed stream) = 0
.
35
gh g�1
,1
L h�1
, distillate/feed = 63
wt
%, Montz
750 structured packing, 45
�
104
◦C
, ambient pressure). The concentrations are
presented in mass fractions. .............................. V
A.20
CO-3, product separation (conditions: 5
.
5
L h�1
, distillate/feed = 82
wt
%,
Montz 750 structured packing, 100
�
210
◦C
, 200
mbar
). The concentrations are
presented in mass fractions. .............................. V
A.21 Pure component properties. .............................. W
A.22 Reft of UNIFAC interaction parameters. . . . . . . . . . . . . . . . . . . . . . . . X
A.23
Deviation of model predicted VLE data and experimental VLE data for four
diferent models. The model Reference contains the model predictions from the
literature sources presenting the experimental data. This work contains the model
predictions from the model used for the process simulation in this work. [108]
contains the model predictions from the model published by Bongartz et. al. [108].
[108]
∗
contains the model predictions from the model published by Bongartz et. al.
[108] updated with the interaction parameters from Schmitz et. al. [91], however
still not considering the temperature dependency of the UNIFAC interaction
parameters. ....................................... Y
A.24
Stream table for the MeOH production based on H
2
and CO
2
. The stream
numbering is presented in Figure A.7. The concentrations are presented in mass
fractions. ........................................ AB
A.25
Operation conditions of the heat exchangers used for the MeOH production based
on H2 and CO2. Numbering is presented in Figure A.7. . . . . . . . . . . . . . . AC
A.26
Operation conditions of the distillation column used for the MeOH production
based on H2 and CO2. Numbering is presented in Figure A.7. . . . . . . . . . . . AC
A.27
Operation conditions of the reactor used for the MeOH production based on H
2
and CO2. Numbering is presented in Figure A.7. . . . . . . . . . . . . . . . . . . AC
xx
List of Tables
A.28
Operation conditions of the phase separators used for the MeOH production based
on H2 and CO2. Numbering is presented in Figure A.7. . . . . . . . . . . . . . . AD
A.29
Operation conditions of the compressors used for the MeOH production based on
H2 and CO2. Numbering is presented in Figure A.7. . . . . . . . . . . . . . . . . AD
A.30
Stream table for the FA(aq.) production from MeOH and air. The stream
numbering is presented in Figure A.8. The concentrations are presented in mass
fractions. ........................................ AE
A.31
Operation conditions of the heat exchangers used for the FA(aq.) production
based on MeOH and air. Numbering is presented in Figure A.8. . . . . . . . . . AE
A.32
Operation conditions of the absorber column used for the FA(aq.) production
based on MeOH and air. Numbering is presented in Figure A.8. . . . . . . . . . AE
A.33
Operation conditions of the reactor used for the FA(aq.) production based on
MeOH and air. Numbering is presented in Figure A.8. . . . . . . . . . . . . . . AF
A.34
Operation conditions of the compressor used for the FA(aq.) production based on
MeOH and air. Numbering is presented in Figure A.8. . . . . . . . . . . . . . . AF
A.35
Stream table for the FA concentration based on FA(aq.) solution. The stream
numbering is presented in Figure A.9. The concentrations are presented in mass
fractions. ........................................ AF
A.36
Operation conditions of the heat exchangers used for the FA concentration based
on FA(aq.) solution. Numbering is presented in Figure A.9. . . . . . . . . . . . . AG
A.37
Operation conditions of the distillation column used for the FA concentration
based on FA(aq.) solution. Numbering is presented in Figure A.9. . . . . . . . . AH
A.38
Operation conditions of the evaporators used for the FA concentration based on
FA(aq.) solution. Numbering is presented in Figure A.9. . . . . . . . . . . . . . AH
A.39
Stream table for the production of OME
3-5
from MeOH and concentrated FA(aq.)
solution. The stream numbering is presented in Figure A.10. The concentrations
are presented in mass fractions. ............................ AI
A.40
Operation conditions of the heat exchangers used for the production of OME
3-5
from MeOH and concentrated FA(aq.) solution. Numbering is presented in Figure
A.10............................................ AI
A.41
Operation conditions of the distillation columns used for the production of OME
3-5
from MeOH and concentrated FA(aq.) solution. Numbering is presented in Figure
A.10............................................ AJ
A.42
Operation conditions of the reactor used for the production of OME
3-5
from MeOH
and concentrated FA(aq.) solution. Numbering is presented in Figure A.10. . . . AJ
A.43 Stream table for the combustion of the purge streams. The stream numbering is
presented in Figure A.11. The concentrations are presented in mass fractions. . . AK
A.44
Operation conditions of the heat exchangers used for the combustion of the purge
streams. Numbering is presented in Figure A.11. .................. AK
A.45
Operation conditions of the reactor used for the combustion of the purge streams.
Numbering is presented in Figure A.11. ....................... AL
xxi
List of tables
A.46
Operation conditions of the compressor used for the combustion of the purge
streams. Numbering is presented in Figure A.11. .................. AL
A.47
Energy demand for H
2
O electrolysis and various CO
2
capture techniques [109].
CPS, CO
2
from point sources; PCC, postcombustion capture; DAC, direct air
capture. ......................................... AL
xxii
1 Introduction
1.1 Context and Background
Anthropogenic climate change already causes devastating consequences all over the world with an
increasing mean temperature, rising sea levels, extreme weather and others events, as illustrated
in Figure 1.1 by [5].
Figure 1.1: Impact of the anthropogenic climate change at 1.5 and 2 ◦C of warming by [5].
To mitigate and fnally reverse the accelerating consequences our mindset, behavior and actions
need to shift from a fossil-based to a carbon neutral and fnally carbon negative as well as circular
economy. The biggest shares of the global CO
2
emissions in 2020 were produced in the energy
1
1 Introduction
sector, namely, electricity and heat (40 %), industry (25 %), transport (21 %) and buildings
(9 %) [6]. Fossil fuels mainly drive the energy economy sectors in a global context, with a total
share of 80 % of the total energy supply (oil 29 %, coal 27 %, natural gas 24 %) [7]. From the
six main technological avenues addressed by the IRENA world energy transitions outlook 2022
[8] and in the net zero emissions by 2050 Scenario (NZE) of the International Energy Agency
(IEA) [9] one major part of the solution is the comprehensive use of renewable energy resources
providing sustainable heat and power that can be integrated into diferent sectors by storing
the energy in various energy carriers [6]. The power-to-X (PtX) concept enables the chemical
storage of renewable electricity by the conversion of H
2
O to H
2
via H
2
O electrolysis. Combined
with CO
2
or N
2
, H
2
can be converted to sustainable energy carriers and chemicals which can be
used for the hard to electrify applications such as the chemical and steel industries but also for
seasonal and large-scale energy storage. Furthermore, for several transportation modes such as
aviation, shipping, heavy duty and other of-road machines, dense liquid fuels will be further
needed. Figure 1.2 illustrats a PtX value chain from sustainable feedstocks by the Fraunhofer
Institute for Solar Energy Systems ISE.
Figure 1.2:
The PtX value chain from sustainable feedstocks via efcient conversion processes
towards advanced PtX products. ©Fraunhofer ISE [10]
A meta-study by Fraunhofer ISE concluded that no energy transformation targets would be met
without the high penetration of H
2
and PtX products in the global energy economy [11, 12].
An average of 550 to more than 1000
TWh
of PtX products should be imported by 2050 in
Germany to fulfll the CO
2
reduction targets of 90 % relative to 1990. One important sustainable
energy carrier is methanol which can be used as a fuel directly or upgraded to other energy
carriers and chemicals, such as oxymethylene dimethyl ethers (OME). OME show promising
fuel and physical properties for a wide range of possible applications such as solvents or diesel
fuel additives or substitutes, which are discussed in the following section 2.1. However, despite
promising properties, a feasible large-scale production of OME is still under investigation.
Various process concepts were proposed for the production of OME from methanol (MeOH)
which are discussed and compared in detail in the following section 4. One of the main challenges
of these processes in terms of technical feasibility and energy demand is the separation of the
by-product H
2
O from the target OME fraction. This H
2
O is formed in various synthesis steps
from MeOH to OME. The OME production processes discussed in the literature use diferent
2
1 Introduction
techniques to separate H
2
O [1, 13–23]. These have diferent advantages and disadvantages partly
coupled with methods which are still in an early phase of investigation and demonstrations. In
the present work, a new process concept is proposed which solves the challenging H
2
O separation
relying on the state-of-the-art reactive distillation technique. This makes the COMET process a
technically feasible process for large production capacities. Besides the comparison of the process
performance to various alternative OME
3-5
production processes, the main process units were
demonstrated and the results are presented as a validation for the feasibility of this new process
concept.
Furthermore, as the heart of the production process, a suitable catalyst is required to selectively
form the target OME product mixture. The synthesis of OME was investigated for various
catalyst systems, feed mixtures and under diferent reaction conditions [24–33]. Besides the
optimization of specifc catalyst properties, some already available commercial catalysts led
to promising results. However, due to the application of diferent feed mixture and reaction
conditions a direct comparison of the results is complicated. Furthermore, besides the conversion,
selectivity, activity, reaction kinetics and side product formation, the thermal stability of the
synthesis product is an important parameter for the process design but was barely mentioned in
the literature. This was investigated in the frame of this work with the result of a recommendation
for a few especially suitable commercial catalysts.
In addition to a suitable catalyst, the process’s overall energy efciency is a deciding factor for
production processes, especially considering the life cycle impact of a PtX product substituting a
fossil-based product. This efciency evaluation depends on the evaluation’s assumptions and,
signifcantly, on the system boundaries. In the context of the transformation and electrifcation
of industrial processes, the industrial high temperature heat pump (HTHP) is a main lever
with signifcant potential to enhance the overall process energy efciency. Around 43 % of the
global energy demand is used as heat in industrial applications. Large amounts of this heat
are exiting the processes as exhaust gases, liquid streams, and cooling water with high shares
of low temperature heat [34]. When upgraded using HTHP, this low temperature heat is a
valuable resource and product in the context of PtX processes. Various studies consider the
determination of the energy efciency of various PtX products and increase the overall energy
efciency of the processes by identifying strategies aiming to reduce the energy demand. In this
work, the energy efciency is additionally enhanced not only by considering the PtX molecule
as the main product of the processes, but also by utilizing the energy streams which cannot be
integrated into the process itself usually addressed as waste streams, due to the low temperature
level. However, the transfer and integration to additional processes which require this low
temperature level reduces the overall energy demand and, therefore, increases the overall energy
efciency. This approach is applied on the expanded OME system boundary and the poten-
tial of the presented technological approach is compared with OME production process efciencies.
3
1 Introduction
1.2 Scope and Objectives of this thesis
The investigations of this cumulative thesis comprise the development of a novel process tackling
the technical challenge of H
2
O separation in the OME value chain relying on reactive distillation
and allowing the industrial large-scale production of OME
3-5
. The COMET process concept
producing OME from MeOH and aqueous FA (FA(aq.)) solutions, was simulated and evaluated
in terms of energy and material efciency identifying key performance indicators and the technical
feasibility. These were compared with alternative processes. Complementary, the experimental
demonstration of the main COMET process units is introduced refecting the succesful production
of a standard compliant OME3-5 product.
Furthermore, another major objective is a consistent comparison of commercial heterogeneous
catalysts for the OME synthesis to identify suitable catalysts available for industrial large-scale
applications. A secondary objective hereby is the investigation of the thermal stability of the
OME synthesis products. This indicates if the synthesis product can directly be separated in a
cascade of distillation columns or if a neutralization step should be considered after the synthesis
to sufciently improve the thermal stability of the mixture and prevent reverse reactions towards
shorter chain OME. This is an important but little discussed aspect of the process concept
development and industrial realization of the OME3-5 production.
Besides the COMET process, alternative OME
3-5
production processes were simulated and
evaluated, with the objective to compare various production process concepts in terms of process
performance and technical feasibility for an industrial large-scale production of OME
3-5
. Thereby,
the overall energy efciency is a key deciding performance indicator and a new approach is
proposed and evaluated to improve the energy efciency by applying HTHP which upgrade low
temperature excess heat streams.
To summarize, this cumulative thesis seeks to answer the following questions::
• What are the performances of the most relevant OME3-5 production process concepts?
•
Which OME
3-5
production process concept is already feasible for demonstration and
scale-up?
•
What are the main hurdles for OME
3-5
production process concepts which hamper their
demonstration and scale-up?
• Which commercial catalysts are suitable for a large-scale OME3-5 production?
• How can the low overall energy efciency of various PtX processes be improved?
4
1 Introduction
Within the framework of this dissertation the results presented in the following chapters and
sections were published in the following four publications:
•
[1] Franz Mantei, Ramy E. Ali, Cornelie Baensch, Simon Voelker, Philipp Haltenort,
Jakob Burger, Ralph-Uwe Dietrich, Niklas von der Assen, Achim Schaadt, Jörg Sauer
and Ouda Salem "Techno-economic assessment and carbon footprint of processes for the
large-scale production of oxymethylene dimethyl ethers from carbon dioxide and hydrogen".
In Sustainable Energy Fuels 6.3 (2022), pp. 528-549. DOI: 10.1039/D1SE01270C
•
[2] Franz Mantei, Matthias Kraume, and Ouda Salem. "Improved Energy Efciency of
Power-to-X Processes Using Heat Pumps Towards Mobility Sector Defossilization". In:
Chemie Ingenieur Technik (2022). DOI: 10.1002/cite.202200118
• [3] Franz Mantei, Sebastian Kopp, Anna Holfelder, Elisa Flad, Daniela Kloeter, Matthias
Kraume and Ouda Salem "Suitable commercial catalysts for the synthesis of oxymethylene
dimethyl ethers". In Reaction Chemistry & Engineering 8 (2023). DOI: 10.1039/D2RE00508E
•
[4] Franz Mantei, Christian Schwarz, Ali Elwalily, Florian Fuchs, Andrew Pounder, Hendrik
Stein, Matthias Kraume and Ouda Salem "A novel process towards the industrial realization
of large-scale oxymethylene dimethyl ether production - COMET". In Reaction Chemistry
& Engineering (2023). DOI: 10.1039/D3RE00147D
A schematical overview of their chronology is illustrated in Figure 1.3. These four publications
are attached in the appendix.
Figure 1.3: Schematical overview of the chronology of the publications this thesis is based on.
In addition, the results were partly published in the following conference contributions:
•
F. Mantei, L. Theiss, O. Salem, A. Schaadt. "The Power-to-Liquid Concept: A Novel
process for the production of (Poly-) Oxymethylene Dimethyl Ether (OME)". 12th European
Congress of Chemical Engineering (2019), Oral presentation.
•
F. Mantei, L. Theiss, A. Schaadt, O. Salem. "(Poly-) Oxymethylendimethylether (OME)
Prozesssimulation mit Aspen Plus
®
". Jahrestrefen der ProcessNet-Fachgruppe Energiever-
fahrenstechnik (2020), Oral presentation.
5
1 Introduction
•
F. Mantei, O. Salem, R. Ali, A. Schaadt. "Prozesssimulation des komplexen Gemischver-
haltens von Oxymethylendimethylether (OME)". 10. ProcessNet-Jahrestagung und 34.
DECHEMA-Jahrestagung der Biotechnologen 2020 (2020), Poster contribution.
•
F. Mantei, R. Ali, O. Salem, A. Schaadt. "Oxymethylenether (OME): Prozesssimulation und
-auswertung". Jahrestrefen der ProcessNet-Fachgruppe Energieverfahrenstechnik (2021),
Poster contribution.
•
F. Mantei, R. Ali, O. Salem, A. Schaadt. "Oxymethylene ethers (OME): Processsimula-
tion and evaluation". 13th European Congress of Chemical Engineering (2021), Poster
contribution.
•
D. Kloeters, F. Mantei, O. Salem, A. Schaadt. "Oxymethylenether (OME): Von der Synthese
zur Destillation". Jahrestrefen der ProcessNet-Fachgruppe Energieverfahrenstechnik (2022),
Poster contribution.
6
2 State of the Art
2.1 OME Properties and Applications
The chemical structure of OME, H
3
C-(O-CH
2
)
n
-O-CH
3
with the chain length
n ≥
1 results in
diferent physical and chemical properties, depending on the chain length
n
. Table 2.1 shows
selected properties of OME1-6 as well as diesel fuel and hydrogenated vegetable oil (HVO).
Table 2.1: Properties of OME1-6, diesel fuel and HVO [35–40].
OME
1
OME
2
OME
3
OME
4
OME
5
OME
6
OME
3-5
EN 590
EN15940
diesel HVO
CAS number
109-87-5 628- 13353- 13352- 13352- 13352- - - -
90-0 3-2 75-5 76- 6 77-7
Molecular weight in
g mol�1
76.1 106.1 136.1 166.2 196.2 226.2 158-166 - -
Boiling point in ◦
C
42 105 158 202 243 273
≥
158 180-370 293.1
Melting point in ◦C
-105 -70 -43 -10 18 38 ---
Density at 15
◦C
in
kg l�1
0.86 0.98 1.03 1.07 1.11 1.14 1.045-
1.07
0.82-
0.845
0.765-
0.80
Kinematic viscosity at
40 ◦C in mm
2 s
�1
0.36
(at 25
◦
C
)
0.67 0.86 1.32 1.93 - 0.9-1.5 2.0-4.5 2.0-4.5
Lower heating value in
M J kg�1
22.4 20.6 19.6 18.9 18.2 17.7 19.0 42.0 43.66
Cetane number
29 68 72 84 93 -
≥
65
≥
51
≥
70
OME are chemically stable in neutral and alkaline environment and non-toxic. Noticable are
increasing boiling and melting points with an increasing chain length, with OME
1
evaporating
already at 42
◦C
and OME
≥ 5
melting just above 18
◦C
. Furthermore, density and viscosity are
increasing with an increasing chain length, but the lower heating value (LHV) is decreasing due
to an increasing amount of oxygen inside the molecular structure. The cetane number is also
increasing with an increasing chain length.
Diferent physical and chemical properties lead to various suitable applications, of which the most
prominent is the application as a fuel. OME
3-5
show similar fuel properties to diesel fuel, a good
solubility with diesel fuel and benefcial combustion behavior [37, 41]. This makes them attractive
as a clean diesel fuel alternative or additive. The high amount of molecular bound oxygen and
absence of C-C bonds enables a quasi soot-free combustion, which can be utilized to reduce
NO
x
emissions. Therefore, the NO
x
and soot emission trade-of of diesel fuel is avoided [39, 42].
Furthermore, by applying OME as a blend with diesel fuel, already low concentrations of OME
can lead to signifcant NO
x
and soot emission reductions for both heavy duty and passenger cars
[41, 43–56]. Under certain conditions, even a mixture of 10
vol
% OME with diesel fuel can lead
to a signifcant reduction in NO
x
and soot emissions [44]. One drawback in comparison to diesel
fuel are the smaller lower heating values of OME of approximately 19
MJ kg�1
in comparison to
42
MJ kg�1
for diesel fuel and 43
.
7
MJ kg�1
for HVO, which leads to higher fuel consumptions
for the same driving distance [37, 39, 44, 57–64]. For blends of diesel fuel with OME or neat
OME, the existing infrastructure for the transportation, storage, and distribution can be used
with some small modifcations, such as sealing materials. Polytetrafuoroethylene (PTFE) is a
compatible thermoplastic and perfuoroelastomer (FFKM) a compatible but expensive elastomer.
For the application of neat OME, ethylene propylene diene monomer (EPDM) is a compatible
8
2 State of the Art
and cheap alternative. However, it is not compatible with diesel fuel [37, 60, 65, 66].
Life cycle assessment (LCA) studies show the potential of CO
2
reduction based on neat OME
or blends with fossil or Fischer-Tropsch diesel fuel, if OME are produced sustainably. For a
certain case study using neat OME
3-5
, Hank et al. [67] evaluated that the well-to-wheel (WtW)
greenhouse gas emissions (GHGE) can be reduced by 86%, corresponding to 29
gCO2�eq. km�1
(OME
3-5
) compared to 209
gCO2�eq. km�1
(diesel fuel). Deutz et al. [68] investigated the WtW
LCA for OME
1
and concluded that it has the potential to serve as an almost carbon-neutral
blending component. Replacing 24
wt
% diesel fuel with OME
1
could reduce the global warming
impact by 22 % and the emissions of NO
x
and soot by 43 % and 75 %, respectively. Voelker et
al. [57] concluded that CO
2
emissions can signifcantly reduce by up to 93 % compared to fossil
diesel fuel. They estimated a NOx reduction of 57 % and an almost complete reduction of soot
using OME instead of diesel fuel. However, already small blending rates of OME show a clearly
positive impact on global CO
2
emissions, as well as local NO
x
and soot emissions [57, 68]. With
a worldwide demand of 26.5 million barrels diesel fuel per day [69], already small blending rates
of OME show the need for large-scale production plants.
Suitable compositions of the fnal OME
3-5
product are defned by key properties, such as density,
viscosity, cetane number and fash point, which are standardized in a new fuel standard for OME
(DIN/TS 51699) [39, 40].
Besides the application as a diesel fuel additive or alternative, OME can be used as a fuel for fuel
cells [70–73]. Furthermore, OME can be used as non-toxic and potentially sustainable solvents.
This application was investigated by Zhenova et al. [74]. They reported good solvation properties
similar to 1,4-dioxane and slow peroxide formations in comparison to tetrahydrofuran (THF).
Due to the miscibility gap of OME with H
2
O, OME can be used for aqueous extractions as well.
Moreover, OME showed a good performance in the dissolution of polysterene and the removal
of paints and coatings. Another potential application for OME is their usage as a solvent for
the production of hydrogen peroxide [75]. Schappals et al. [76] investigated the application
of OME as physical absorbens for CO
2
. They reported good solubilities with lower enthalpies
of absorption in comparison to the currently used absorbents selexol and sulfolane which is
benefcial in terms of energy efciency.
2.2 OME synthesis
Based on methanol (H
3
C-OH, MeOH), various synthesis routes for the production of OME
(H
3
C-O-(CH
2
O)
n
-CH
3
) take place over acid catalysts in the liquid phase at temperatures usually
between 50
�
100
◦C
[25]. For the synthesis of OME, methyl capping group suppliers such as
MeOH, methylal (H
3
C-O-(CH
2
O)
1
-CH
3
, OME
1
) or dimethyl ether (H
3
C-O-CH
3
, DME) react
with a formaldehyde source (H
2
C-O, FA) such as formalin, paraformaldehyde (HO-(CH
2
O)
n
-H
with n = 8-100, pFA), trioxane ((CH
2
O)
3
, TRI), or monomeric FA through an initiation, growth,
and termination mechanism, as described by Baranowski et al. [77]. This leads to several
simultaneous and successive reactions and the formation of intermediate and side products.
In a solution of MeOH and H
2
O, FA is bound in poly(oxymethylene) hemiformals (HO-(CH
2
O)
n
-
9
2 State of the Art
CH
3
with
n
= 1
�
10, HF
n
) following eqn. 2.1 and 2.2 and poly(oxymethylene) glycols (HO-
(CH
2
O)
n
-H with
n
= 1
�
10, MG
n
) following eqn. 2.3 and 2.4. These reactions are fast, even in
absence of a catalyst [78–80]. In solutions with MeOH and H
2
O the amount of monomeric FA is
very small towards chemical equilibrium [80].
CH2O + CH3OH ⇄ HO(CH2O)1CH3 (2.1)
CH2O + HO(CH2O)n�1CH3 ⇄ HO(CH2O)nCH3; n = 2 � 10 (2.2)
CH2O + H2O ⇄ HO(CH2O)1H (2.3)
CH2O + HO(CH2O)n�1H ⇄ HO(CH2O)nH; n = 2 � 10 (2.4)
HF
n
and MG
n
are unstable due to the fast reactions at changing compositions and conditions.
However, their formation could be experimentally investigated and quantifed by the group of
Hasse et al. [78, 81–83] using NMR techniques.
In an acidic environment the acetalization reaction of MeOH and HF
1
towards OME
1
proceeds
as follows [79]:
+
⇄
H
The chain propagation of OME proceeds following an acetalization mechanism with a sequential
growth mechanism as described by eqn. 2.6 and 2.7, respectively [80].
HO(CH2O)1CH3 + CH3OH CH3O(CH2O)1CH3 + H2O (2.5)
+
⇄
H
CH2O + CH3O(CH2O)n�1CH3
HO(CH2O)nCH3 + CH3OH CH3O(CH2O)nCH3 + H2O; n = 2 � 10 (2.6)
+
⇄
H
In addition, fast transacetalization reactions as described by eqn. 2.8 support the chain dis-
tribution of the OME molecules which can be described by a Schulz-Flory distribution [84,
85].
CH3O(CH2O)nCH3; n = 2 � 10 (2.7)
+
⇄
+
⇄
+
⇄
+
⇄
H
(2.8)
The main side products formed in the OME synthesis are methyl formate (HCOOCH
3
, MEFO),
formic acid (HCOOH, FOAC), DME and TRI [28, 86]. TRI is also used as a feedstock for the
OME synthesis and can be formed following diferent mechanisms as described by eqn. 2.9-2.11
[28, 87].
H
H
H
CH3O(CH2O)nCH3+CH3O(CH2O)mCH3O(CH2O)n+kCH3+CH3O(CH2O)m�kCH3
CH3
3CH2O (CH2O)3 (2.9)
HO(CH2O)3H (CH2O)3 + H2O (2.10)
CH3O(CH2O)4CH3 (CH2O)3 + CH3O(CH2O)1CH3 (2.11)
10
2 State of the Art
MEFO can be formed as a combination of two FA molecules following the Tishchenko reaction, as
described by eqn. 2.12 [26] or from FOAC and MeOH via a reversible esterifcation as described
by eqn. 2.13 [28].
2CH2O H
→
+
HCOOCH3 (2.12)
H+
HCOOH + CH3OH ⇄HCOOCH3 + H2O (2.13)
FOAC can also be formed from MEFO as described by eqn. 2.13 or as a combination of two
FA molecules and H
2
O in presence of an acidic or alkaline catalyst, following the Cannizzaro
reaction as described by eqn. 2.14 [28].
H+/OH�
2CH2O + H2O → HCOOH + CH3OH (2.14)
DME can be formed from two MeOH molecules or in a backwards reaction from OME
1
, as
described by eqn. 2.15 and 2.16 [26, 88].
H+
2CH3OH ⇄CH3OCH3 + H2O (2.15)
H+
CH3O(CH2O)1CH3 ⇄CH3OCH3 + CH2O (2.16)
The syntheses towards OME can be divided into aqueous reaction systems comprising the
presence of H
2
O in the reaction mixture, and anhydrous systems without the formation of H
2
O
[89]. H
2
O is formed if MeOH is directly used for the OME synthesis as described by eqn. 2.5
and 2.6. Moreover, H
2
O can enter the synthesis together with the FA source such as formalin
or pFA. Its presence leads to the formation of the side products HF and MG and reduces the
selectivity towards OME
3-5
[27, 90]. Furthermore, the product purifcation is more complex
due to several azeotropes, complex vapor-liquid-liquid equilibria (VLLE), challenges regarding
FA solidifcation and the separation of H
2
O from the process [18, 91–93]. On the other hand,
anhydrous reaction systems lead to a much simpler product purifcation. However, in this
case H
2
O needs to be separated from the feedstock before entering the OME reactor, which is
especially energy-intensive for the production of the reactant TRI [87, 94].
For the investigation of the OME synthesis, two reaction systems were evaluated with MeOH
and pFA as a typical aqueous system and OME1 and TRI as a typical anhydrous system.
2.3 Catalyst systems for the synthesis of OME
OME are synthesized in an acidic environment with Brønsted and Lewis acid sites activating the
synthesis. Lewis acid sides are active for the decomposition of pFA as described by eqn. 2.4,
acetalization and chain propagation of OME as described by eqn. 2.5-2.7. Brønsted acid sites
are active for all steps of the OME synthesis, including the ring-opening of TRI [77]. Various
liquid and solid catalyst systems were already applied to the OME synthesis including acidic ion
exchange resins (IER), zeolites and ionic liquids.
Oestreich et al. [24] investigated the OME synthesis from a mixture of MeOH and pFA over the
11
2 State of the Art
IER Dowex50Wx2, Dowex50Wx4, Dowex50Wx8, Amberlyst
®
36 (A36) and IR-120 and ground
zeolites H-MFI 90, H-BEA 25, CBV 720, H-MFI 400 and H-MOR 30 in a batch autoclave at
80
◦C
. They compared the activity of the catalysts by determining the time after which 9 wt%
of OME
2
were obtained. Their results show that the Dowex catalysts had by far the highest
activity, followed by the zeolites H-BEA 25 and H-MFI 90. A36 and IR 120 showed a lower
activity and H-MFI 400 and H-MOR 30 did not reach the required OME
2
concentration after
100 min. Regarding the side product formation, TRI and MEFO were detected far below 1 wt%
for all IER but are pronounced for the zeolites with concentrations higher 1 wt% at 80
◦C
and
longer retention times.
Lautenschütz [25] investigated the OME synthesis from OME
1
and TRI over the IER A15, A16
and A36, ground zeolites H-BEA 25, H-BEA 150, H-BEA 300, H-FER 20, H-MFI 27, H-MFI 90,
H-MFI 240, H-MOR 30 and H-FAU 30 and
γ
-Al
2
O
3
in a batch autoclave and fxed bed reactor
at 30
�
100
◦C
. At 40
◦C
, A15 was more active than A36 and A16. Furthermore, Lautenschütz
reported that A15, A16 and A36 were more active than A46, comparing his results to the results
from Burger et al. [26]. Regarding the zeolites, H-BEA 25, H-BEA 150, H-BEA 300 and H-FAU
30 were active for the OME synthesis while the other zeolites only reached low conversions at the
same retention time. Moreover, in a similar form and particle size the IER were still more active
than the zeolites for both powder form and grain shape. However, the diference between grain
shape to powder form led to a much higher activity for A36 (factor 16) and comparatively small
improvements for H-BEA 25 (factor 3).
γ
-Al
2
O
3
was not active for the OME synthesis from
OME
1
and TRI. For the reactants OME
1
and pFA similar results were obtained with a higher
activity for the IER followed by the BEA zeolites. However, the activity reduced signifcantly
in comparison to the OME
1
and TRI feed mixture, due to the presence of H
2
O which leads to
the formation of several side products. For the reactants MeOH and pFA as well as MeOH and
TRI the activity reduced for A36 and H-BEA 25, which was stronger pronounced for H-BEA
25. Regarding the side product formation, Lautenschütz [25] reported no detection of MEFO
at 40 ◦C for A15, A16 and A36 but detected MEFO for A36 at temperatures above 60 ◦C. No
observations were reported regarding the TRI side product formation for OME
1
and pFA or
MeOH and pFA.
Burger et al. [26] investigated the OME synthesis from OME
1
and TRI over the IER A36 and
A46 in a batch autoclave at 50
�
80
◦C
. They reported that A36 led to the formation of 1
�
2 wt%
DME and MEFO as side products, while in the tests using A46 no DME or MEFO could be
detected. Schmitz et al. [27] investigated the OME synthesis from MeOH and pFA over A46
in a batch autoclave at 60
�
105
◦C
. They detected MEFO and TRI as side products. TRI
concentrations of up to 2.6 wt% were obtained for feed mixtures with high FA concentrations
and high reaction temperatures. The MEFO concentration did not exceed 0.06 wt% and was
mainly below the detection limit. Voggenreiter et al. [28] investigated the side product formation
over A46 for mixtures of FA, MeOH, H
2
O and OME
1
and published a kinetic model for the
formation of MEFO, TRI and FOAC. High MEFO concentrations above 1 wt% were only detected
at temperatures greater or equal to 85
◦C
and long residence times far after the equilibrium
composition of the OME was obtained. The FOAC concentration was mainly a bit lower than
the MEFO concentration but followed a similar behavior. The TRI concentration never exceeded
12
2 State of the Art
1 wt% and was limited by the equilibrium composition.
Zheng et al. [29] investigated the OME synthesis from OME
1
and pFA over the IER NKC-9,
D001-CC and D72 in a batch autoclave at 20
�
80
◦C
. They reported a high activity for NKC-9.
No formation of side products was reported.
Wu et al. [30] investigated the OME synthesis from OME
1
and TRI over the zeolite H-MFI
with various Si/Al ratios in a batch autoclave at 120
◦C
. They reported an increasing OME
1
conversion and decreasing TRI conversion for higher Si/Al ratios and a decreasing formation of
MEFO, while MeOH and FA concentrations were increasing. Above the molar Si/Al ratio of 580
the MEFO concentration decreased below 1 wt%.
Wang et al. [31] investigated the OME synthesis from OME
1
and TRI over various homogeneous
and heterogeneous catalysts, including the zeolites H-FAU, H-MFI and the IER A15, D002, D009
and CT175 in a batch autoclave at 90
◦C
. They reported a low activity for the zeolites and a
high activity and selectivity for the IER, especially for CT175. However, they did not report the
detection of side products but mentioned the formation of pFA for very low OME
1
to TRI ratios.
Fink et al. [32] investigated the OME synthesis from MeOH and pFA over the zeolites H-BEA 13,
H-BEA 18, H-BEA 81, H-FAU 3, H-FAU 15, H-FAU 35, H-FAU 49, H-MFI 14, H-MFI 34, H-MFI
114, H-MFI 4716, H-MOR 6, H-MOR 10, H-MOR 16 in a batch autoclave at 65
◦C
. They did
not investigate the performance of IER catalysts since in preliminary experiments, the leaching
of SO
3
H groups was detected, leading to signifcant sulfur contents in the synthesis product. The
catalysts with the highest activity for the OME synthesis were H-BEA 81, H-BEA 18, H-MFI 34
and H-FAU 35. The H-MOR zeolites showed far lower activities than all other zeolites except
H-FAU 3 and H-MFI 4716 which showed very low and no activity, respectively. Additionally,
they reported a minor formation of MEFO for all active catalysts with concentrations of about
0.1 wt% for most catalysts and 0.3 � 0.4 wt% for the H-MFI catalysts.
Endres et al. [33] investigated the microwave-assisted OME synthesis from OME
1
and TRI over
the IER A15, A36, Dowex50Wx2, Dowex50Wx4 and Nafon in microwave vials at 25
�
100
◦C
.
Their results show a higher activity for A15 than Nafon but a far lower formation of MEFO for
Nafon than for all IER, especially at 40
◦C
. The observation of MeOH and FA in the samples
was not reported.
This work focuses on the heterogeneous OME synthesis from commercial catalysts which could
be used in industrial applications. The catalysts used for this evaluation are listed in Table 2.2
and were selected based on the reported performances regarding the OME synthesis from the
abovementioned investigations.
2.4 OME3-5 production processes
OME production processes have been investigated intensively since the early work by DuPont in
the middle of the 20th century on the production of longer chain OME [25]. Since the 1990s, short
chain OME have been recognized as interesting diesel blends or substitutes [102]. Subsequently,
intensive research eforts have been focused on engine testing on one hand, and production
processes on the other hand, led by the Fordmotor company and Eni SpA [103, 104]. At the
beginning of the 21st century, fundamental developments led by BASF and BP established
13
2 State of the Art
Table 2.2: Catalysts for the OME synthesis from MeOH-pFA as well as OME1-TRI [1]
Catalyst
Type Form Surface area
in m
2
g
-1
Acid capacity
in meq g
-1
Si/Al
ratio
T
max
in
◦C
Amberlyst® 15 (A15)
IER
Spherical
53 [95] 4.7 [95] - 120 [95]
Amberlyst® 36 (A36)
IER
Spherical
33 [96] 5.4 [96] - 150 [96]
Amberlyst® 46 (A46)
IER
Spherical
75 [14] 0.8-1.3 [97] - 120 [25]
Dowex
®
50WX2 IER
Spherical
- 0.81 [98] - 150 [98]
(Dowex)
H-BEA 25
Beta
Cylindrical
>500 [99] - 25 [99] >200 [98]
zeolite
H-MFI 90
Pentasil
zeolite
Cylindrical
>300 [100] -
30 [100]
>200 [98]
Nafon™NR40
(Nafon)
Perfuoro-
sulfonic
Spherical
0.001
a
1.0 [101] - 200 [101]
acid resin
a Assumption: diameter of 3 mm, density of 2 g cm-3, no pores.
production processes for OME on research and pilot scales. Most of the following contributions
on the process side were led by Chinese research and industrial groups, especially the important
work by China Petroleum & Chemical Corporation SINOPEC [25, 105]. An overview of the
publications and the patents, together with the research activities in Germany and worldwide
considering OME is given in the review work by Hackbarth et al. [89] and Lautenschütz [25],
elaborating the intensity of research in this feld. Currently, some OME production plants are in
operation or under construction in China with production capacities of 10
�
400
kt a�1
mostly
based on the feedstock OME
1
or MeOH and pFA or TRI [89, 106]. However, information are
scarce about their performance, the quality and composition of the fnal OME product and the
long-term operation.
2.5 H2O separation from the production of OME
One of the main challenging and energy-intensive process steps is the separation of the by-product
H
2
O from the production chain towards OME
3-5
. Considering a sustainable production of OME
3-5
based on MeOH produced from H
2
and CO
2
, H
2
O is formed in the MeOH synthesis, synthesis of
intermediate products such as FA(aq.), DME and OME
1
and the OME
n
synthesis, as illustrated
in Figure 2.1.
In the MeOH synthesis from H
2
and CO
2
, the OME
1
synthesis and the DME synthesis, H
2
O
is a by-product and separated using distillation columns [1, 107, 108]. In the partial oxidation
of MeOH towards FA, as described by eqn. 4.4, H
2
O is formed as a by-product and used as a
washing liquid in the absorber column. Downstream, H
2
O is partly separated in a concentration
step [1, 109]. However, FA(aq.) is concentrated but not completely separated from H
2
O using
evaporation techniques. Therefore, H
2
O is introduced into the TRI synthesis and separated in
an energy-intensive cascade of distillation columns [87, 94]. Only in the anhydrous FA synthesis,
which is still in its very early stages, no H
2
O is present [110]. Regarding the synthesis of OME
≥ 2
,
14
2 State of the Art
Figure 2.1:
H
2
O separation from a sustainable production of OME
3-5
based on H
2
and CO
2
[4].
a combination of the methyl group supplier DME or OME
1
and oxymethylene groups suppliers
TRI or monomeric FA does not lead to the formation of H
2
O as a by-product, as described by
eqn. 2.7 and 2.9. This simplifes the fnal product purifcation.
However, starting from the cheaper and established reactant FA(aq.), H
2
O will always be present
in the OME
3-5
sub-process and needs to be separated from the loop to circumvent accumulation.
Due to a complex phase behavior of the synthesis product mixture containing mainly FA, H
2
O,
MeOH, OME
1-10
, HF and MG with several azeotropes and close boiling points, H
2
O cannot
be separated individually simply via distillation. The separation of H
2
O from the loop is still
a major challenge regarding the implementation of a potentially cheaper and simpler aqueous
OME3-5 production process.
2.5.1 Extraction
Using an extractant for the separation of H
2
O from the OME synthesis mixture separates the
mixture into two phases, one organic phase mainly containing OME and the extractant and one
aqueous phase mainly containing H
2
O, FA and MG. Downstream to the extraction the organic
phase can be separated and purifed using distillation columns. The extractant is also separated
and recycled back to the extraction. Various extractants were investigated in the literature
showing that toluene, p-xylene and n-heptane enable promising liquid-liquid separation behaviors
between OME and H2O, FA and MG [111–117].
Results of Li et al. [112] show the separation of the OME synthesis product using toluene. About
70 % of OME are separated in the organic phase and only 14 % of FA and H
2
O migrate in the
organic phase, as indicated by the split fraction. However, the organic phase mainly consists of
toluene, which needs to be separated to be recycled. Furthermore, FA and H
2
O still represent a
large proportion of the organic phase and the aqueous phase still contains a large proportion of
OME. A graphical illustration of the results of Li et al. [112] is presented Figure 2.2.
In addition to the extraction method illustrated in Figure 2.2, Oestreich et al. [118] investigated
the extraction of OME from H
2
O, MeOH, FA and TRI using hydrocarbons, i.e. n-dodecane,
15
2 State of the Art
Figure 2.2:
H
2
O separation from the OME synthesis product using toluene as extractant (condi-
tions: (OME synthesis product)/toluene = 0
.
667
gg
�1
, 25
◦C
, batch) by Li et al.
[4, 112].
diesel and HVO as extractants. They proposed to use the hydrocarbons already during the
synthesis to gain a product phase containing OME
n
with hydrocarbons and an aqueous phase
containing MeOH, FA, H
2
O and the catalyst. For the application as a fuel OME
1
should be
separated to increase the fash point which will also separate most of the H
2
O content in the
product phase. Their analysis showed that the remaining mixture of HVO or diesel fuel with
about 7 wt% OME2-10 complies to current fuel standards to a large extend.
2.5.2 Adsorption
Schmitz et al. [18] investigated the adsorption of H
2
O from a mixture containing FA, H
2
O,
MeOH and OME
1-4
using zeolite 3A. Their results show that zeolite 3A has a good selectivity
for H
2
O with only small amounts of FA and MeOH being separated from the feed mixture. A
graphical illustration of the results is presented in Figure 2.3. Ferre et al. [93] investigated the
adsorption of H
2
O using zeolite 3A from binary and ternary mixtures with MeOH and FA. Their
results show that an increasing amount of FA or MeOH in the feed mixture leads to an increased
adsorption of these components. However, the selectivity for H2O is still far higher.
Regarding the separation of H
2
O from an OME
3-5
production process the adsorption has the
advantage of selectively separating H
2
O from the loop, which enables the recycle of all other
components to the OME synthesis. Due to the reaction network between H
2
O and FA as
described by eqn. 2.3 and 2.4 not only the monomeric H
2
O is separated, but also H
2
O from MG
n
.
Therefore, a signifcant reduction of the overall H
2
O content can be achieved. However, without
H
2
O, FA from MG
n
remains in the mixture and either bounds with HF
n-1
to long chain HF
n
, or
with MG
n-1
to long chain MG
n
or it remains in monomeric form. Either way, it increases the risk
16
2 State of the Art
Figure 2.3:
H
2
O separation from the OME synthesis product via adsorption using zeolite 3A
(conditions: (OME mixture)/(zeolite 3A) = 2
.
0
g g�1
, 25
◦C
, batch) by Schmitz et
al. [4, 18].
of local precipitations and, therefore, deactivation of the adsorbents. Therefore, a regeneration
might be necessary. To reduce the risk of precipitation the temperature can be lifted, or the
remaining H
2
O content can be increased. The latter would, however, decrease the yield of
OME
3-5
in the OME synthesis and, therefore, increase the recycle streams and heat demand
for separation. A suitable remaining H
2
O content should be experimentally investigated and
confrmed by long-term stability tests with alternating sequences of adsorption and regeneration.
Furthermore, the scale-up potential should be investigated to ensure its feasibility for large-scale
production plants.
Regarding the heat demand for the separation of H
2
O via adsorption, Schemme et al. [119,
120] estimated that 2
.
1
kWh kg�1
are required. Their estimations are based on the results
H2O
from Schmitz et al. [18, 121] and assume that the adsorbents are heated up from 25
◦C
to 235
◦C
for the regeneration using high pressure steam. Furthermore, it was assumed that
the heat demand is mainly based on the heat of adsorption and the heat capacity of the adsorbents.
2.5.3 Membrane
Schmitz et al. [92] tested two zeolite membranes type NaA and type T from Mitsui & Co. as well
as three PVA-based polymer membranes PERVAP 4100, PERVAP 4101 and PERVAP 4102 from
DeltaMem AG for a mixture containing FA, H
2
O, MeOH, OME
1
and OME
2
. Their results show
that the zeolite membranes and the PERVAP 4102 were not suitable for the separation task,
while the PVA-based polymer membranes PERVAP 4100 and PERVAP 4101 could separate H
2
O
with a high selectivity, also for the repeated experiment with the used membrane. A graphical
illustration of the results by Schmitz et al. [92] using the PERVAP 4100 membrane is presented
in Figure 2.4.
However, Ferre et al. [122] reported the application of a diferent membrane from DBI Gas und
Umwelttechnik. The long-term stability of the membrane, selectivities in the reaction mixture
17
2 State of the Art
Figure 2.4:
H
2
O separation from the OME synthesis product using the polymeric membrane
PERVAP 4100 (conditions: 80
◦C
,2
mbar
permeate, 80
lh
�1
) by Schmitz et al. [4,
92].
and the scale-up potential should be further investigated to ensure its feasibility for large-scale
production plants.
The advantages of the membrane for the separation of H
2
O are similar to the advantages of the
adsorption with a high selectivity for H
2
O. However, likewise to the adsorption, this results in
a higher risk for local precipitation. Therefore, a compromise might be necessary between the
long-term stability and the H
2
O concentration of the retentate. A disadvantage of a higher H
2
O
concentration in the retentate is an increase of the recycle streams which results in higher heat
demands for the product purifcation and reduces the overall energy efciency of the process.
Regarding the heat demand for the separation of H
2
O via membranes, Held et al. [109] estimated
that 0
.
7
kWh kg�1
are required. This results from the evaporation of H
2
O after passing through
H2O
the membrane to the reduced pressure of the permeate of 0
.
03
bar
. Held et al. [109] assumed
that no external heat is required but the temperature of the process stream is reduced from
84
◦C
to 36
◦C
. In comparison to the separation of H
2
O via adsorption, the heat demand is
signifcantly lower.
Table 2.3 summarizes the main advantages and main hurdles of the H
2
O separation methods
extraction, adsorption, and membrane.
Table 2.3:
Main advantages and main hurdles of the H
2
O separation methods extraction, adsorp-
tion, and membrane [4].
Method Extraction Adsorption Membrane
H2O selectivity
Energy demand
Long-term operation
Scale-up potential
Low
High
Likely
Likely
High
High
Challenging
Challenging
High
Comparatively low
Challenging
Challenging
18
3 Methods
3.1 Experimental investigation of the OME synthesis
3.1.1 Chemicals and materials
The reactants MeOH (purity
≥
99
.
9 %), granulated pFA (purity 94
.
5
�
95
.
5 %) and OME
1
(purity
≥
99
.
9 %) were purchased from Carl Roth GmbH + Co. KG. TRI (purity
≥
99 %) was purchased
from Sigma-Aldrich Chemie GmbH. OME
2
(purity
≥
98
.
5 %), OME
3
(purity
≥
99 %), OME
4
(purity
≥
98
.
5 %) and OME
5
(purity
≥
98
.
5 %) were used for calibration and were supplied by
ASG Analytik-Service AG. MEFO (purity 97 %, 3 % MeOH) was purchased from Thermo Fisher
GmbH. Anhydrous sodium sulfte (purity
≥
98 %) and sulfuric acid (
C
= 0
.
1
mol L�1
, ±
0
.
2 %)
were purchased from Carl Roth GmbH + Co. KG. The solvent EtOH (ethanol, purity
≥
99
.
9 %),
the indicator thymolphthalein and the internal standard ethyl acetate (EA, purity
≥
99
.
9 %)
were purchased from Carl Roth GmbH + Co. KG. HYDRANAL™-Solvent Oil and -Titrant 5
were purchased from Honeywell International Inc. Fluka. All chemicals were used without further
purifcation.
The catalysts A15, A36 and Dowex were purchased from Sigma-Aldrich Chemie GmbH, Nafon
was purchased from Ion Power GmbH. A46 was provided by INAQUA Vertriebsgesellschaft mbH
and the zeolites H-BEA 25 and H-MFI 90 were provided by Clariant AG. The IER and zeolites
were dried overnight at 20
mbar
and 30
◦C
before use. IER III was purchased from Merck
Chemicals GmbH.
3.1.2 Analytics
The quantitative analysis was performed using an Agilent 7890A gas chromatograph equipped
with a fame ionization detector (GC-FID) to analyze the organic components of the obtained
reaction products. A sample volume of 1
µL
was injected by an Agilent 7693A autosampler onto
a DB-5MS column (
l
= 30
m
,
di
= 0
.
25
mm
, flm thickness = 0
.
5
µm
). He (g) was used as carrier
gas (fow rate: 202
.
5
mL min�1
,
p
= 11
.
154
psi
, split ratio = 200 : 1). The GC inlet temperature
was set at 290
◦C
, the temperature of the oven was programmed as a ramp (
T5min
= 30
◦C
,
Tramp
= 30
◦C min�1
,
T7min
= 270
◦C
,
ttotal
= 20
min
). Calibration of the GC was achieved
using EA as internal standard (
Ai/AEA
=
Ri · wi/wEA
). The components OME
1-5
, MeOH and
TRI were calibrated using pure mixtures. MEFO was calibrated using a 97 % pure mixture and
subtraction of the MeOH content of 3 %. The components OME
6-11
were calibrated based on
extrapolation and relating the peak area ratio per mass fraction ratio to the internal standard as
a linear function of the number of carbon atoms of the OME molecules.
The H
2
O content of the obtained samples was determined by Karl-Fischer titration and the
content of FA was determined using the sodium sulfte method. Both methods quantify the
overall composition, including H
2
O and FA bound in HF and MG. For a consistent data set,
20
3 Methods
the analyzed H
2
O and FA contents were not adjusted. Still, the content of the components
quantifed via GC-FID were normalized by proportional weighing to a sum of 1
g g�1
. The sum
of the overall mass fractions before this adjustment was predominantly between 0
.
85
g g�1
and
1
.
05
g g�1
for all samples. The main challenge was the precise analysis of MeOH due to the
unstable side products HF which contain MeOH and were detected by the GC-FID. This was
partly compensated by adjusting the MeOH calibration.
Due to their fast reaction kinetics of HF and MG as described by the eqn. 2.1-2.4, these molecules
are unstable at changing compositions and conditions. In this work, the true composition and,
therefore, the formation of HF and MG is not considered. For the evaluation and presentation of
the results, the overall composition is used considering the decomposition of HF and MG to their
constituents MeOH, H
2
O and FA. This does not change the mass fraction of OME and other
side products and, therefore, does not infuence the conclusion from the results. Nevertheless, it
strongly simplifed the analysis of the samples.
For the OME
1
-TRI feed mixture the components FA, MeOH, H
2
O and MEFO were considered
as side products. For the MeOH-pFA feed mixture only TRI and MEFO were considered as side
products.
3.1.3 Apparatus
For the synthesis of OME, a high-pressure laboratory autoclave (
Vmax
= 500
mL
,
pmax
= 100
bar
)
was used in combination with an integrated heating jacket (
Tmax
= 300
◦C
) and a magnetic
stirrer from Carl Roth GmbH + Co. KG. The temperature was measured using a NiCr-Ni
thermocouple (K-type; accuracy
±
1
.
5
K
). The pressure was measured using a diaphragm
pressure indicator (accuracy
±
0
.
24
bar
) [123]. The lid of the autoclave was designed and adapted
to the experimental requirements. The sampling line was cooled using a counter current heat
exchanger operated with tap water to avoid evaporation of the samples. In addition, a sintered
stainless-steel flter (pore size = 10
µm
) near the reactor bottom ensured sampling without
catalyst particles. A scheme of the reaction setup is illustrated in Figure 3.1a.
For the distillation of the OME synthesis product mixtures, a micro distillation setup was used
which was heated using an oil bath and a magnetic stirrer inside the 50
mL
two-neck round-
bottom fask was used for mixing. No column was used to enable the distillation of smaller sample
amounts. A Liebig condenser with a thermometer and vacuum nozzle was used to condense the
distillate and collect the product in 10
mL
fasks. The temperature in the bottom product was
measured using a NiCr-Ni thermocouple (K-type; accuracy
±
1
.
5
K
). The oil bath was heated with
a heating plate regulated with a thermocouple inside the oil bath and a magnetic stirrer inside the
oil bath for faster heat distribution. A scheme of the distillation setup is illustrated in Figure 3.1b.
3.1.4 Feed preparation
The MeOH-pFA feed mixtures were prepared by dissolving pFA in MeOH at a ratio of pFA/MeOH =
1
.
5
g g�1
corresponding to the maximum FA solubility in a methanolic solution, similar to the
21
3 Methods
(a) (b)
Figure 3.1:
Autoclave (a) and distillation setup (b) for the investigation of the OME synthesis
and thermal stability of the OME synthesis product [3]. BV, back pressure valve; F,
flter; H, heat exchanger; P, product fask; PI, pressure indicator; S, catalyst chamber;
TI and TIC, temperature indicators; V, valve.
procedure from Oestreich et al. [24]. pFA was dissolved by stirring and heating to 85
◦C
for up to
three days until a clear solution was obtained. A condenser was placed above the round bottom
fask to condense evaporating components. After cooling the mixture to ambient temperature, it
was fltered using a pleated flter (retention range: 5
�
8
µm
) to avoid solid pFA particles in the
feed mixture.
OME
1
-TRI feed mixtures were prepared by dissolving TRI in OME
1
at a ratio of OME
1
/TRI =
2
g g�1
. TRI was dissolved at ambient temperature by stirring for two hours until a clear solution
was obtained. To avoid the evaporation of OME
1
, the round-bottom fask was closed with a lid
and fltration of the feed mixture was omitted.
3.1.5 OME synthesis
A pressure test was performed before each synthesis experiment at 8
bar
. For the OME synthesis,
1
wt
% of catalyst (
mcat
= 3
.
5
g
) was used in comparison to the mass of the feed mixture
(
mfeed
= 350
g
). The catalyst was added above the reactor head in a catalyst chamber. The
prepared feed mixture was added to the reaction chamber without contact to the catalyst. After
the autoclave was completely sealed again, N
2
was added until a pressure of 2
bar
was reached.
The line above the catalyst was flled with N
2
at 8
bar
. The feed mixture was heated up to
60
◦C
inside the autoclave at constant stirring rate. At 60
◦C
the catalyst was added to the
22
3 Methods
feed mixture by opening a ball valve above the reactor lid and N
2
was added until 8
bar
. With
the addition of the catalyst, the synthesis experiment started. During the synthesis, samples
were withdrawn through a cooled sampling line. To avoid contamination inside the sampling
line, it was rinsed by withdrawing 5
mL
of reaction product before withdrawing the sample. The
frst sample (
S
0) was withdrawn from the feed mixture before it was fed to the reactor. The
second sample (
S
1) was withdrawn from the reactor when a temperature of 60
◦C
was reached.
All further samples were withdrawn at progressively longer time intervals after the catalyst was
added to the reaction chamber (
t
= 1
,
5
,
10
,
15
,
20
,
30
,
45
,
60
,
75
,
90
,
120
,
150
,
180
,
240
,
300
min
).
The fnal sample (
S
17) was taken together with the reaction product after 24
h
. After the
withdrawal of the product mixture the autoclave setup was completely cleaned.
The reaction temperature of 60
◦C
was chosen as a compromise between the reaction kinetics of
the MeOH-pFA and the OME
1
-TRI feed mixture, the amount of catalyst and the formation of
side products. In the literature mentioned in section 2.3, the OME synthesis was investigated
at various temperature levels without signifcant changes in the fnal OME distribution but
signifcant changes in the side product formation.
3.1.6 Catalyst performance evaluation
The performance of the catalysts was evaluated based on the conversion
X
of the feed, the molar
fraction of OME
3-5
in the range of products, further described as selectivity
S
, the mass fraction
of OME
3-5
in the synthesis product mixture, further described as yield
Y
, the activity and side
product formation. The conversion of the feed was evaluated with the mass fraction
w
of the
feed before the synthesis at 0 h and after the synthesis at 24 h, as described by eqn. 3.1 [25].
wfeed,0h � wfeed,24h
Xfeed = (3.1)
wfeed,0h
The selectivity of OME
3-5
was evaluated with the mole fraction
x
of OME
3-5
and all the products
quantifed, considering H2O, OME1-10, MEFO and FA and MeOH or TRI, as described by eqn.
3.2.
xOME3�5
SOME3�5 = (3.2)
xP roducts
The yield of OME
3-5
was evaluated as described by eqn. 3.3 with the mass of the products and
educts [25].
mOME3�5
YOME3�5 = (3.3)
mP roducts + mEducts
The activity was evaluated using two indicators. The frst indicator was the termination time
tT ermination
which is the time after which 90 % of the OME
5
concentration after 24
h
was
obtained, as described by eqn. 3.4.
wOME3�5 (tT ermination)=0.9 · wOME5 (24h) (3.4)
tT ermination
was determined via linear interpolation of the progress of the OME
5
concentration.
It indicates when the fnal product formation was approximately reached, and the reaction could
23
3 Methods
be terminated. This is especially important considering the formation of side products like
TRI and MEFO whose concentrations increase with increasing residence time. However, the
linear interpolation leads to increasing errors for high gradients and large time steps between the
samples.
The second indicator is the relation of the yield of OME
3-5
after 30
min
to the yield of OME
3-5
after 24 h, as described by eqn. 3.5.
YOME3�5 (30 min)
Y ∗ = (3.5)
OME3�5 YOME3�5 (24 h)
3.1.7 Distillation of the OME synthesis products
The synthesis products were directly distilled in a micro distillation setup to investigate the
necessity of a product neutralization step before the thermal separation for the target product
purifcation. The thermal stability of the product mixtures thereby determines whether the
synthesis product can be directly purifed in a distillation column or if the process concept needs
to be extended by a neutralization step.
After the micro distillation setup was mounted, the weight of the round-bottom fask was
measured, and 30
�
50
g
of the synthesis product was added to the round-bottom fask. Then,
the oil bath was heated stepwise from 60
◦C
to 100
◦C
. Thereby, the temperature was increased
after the distillate fow stopped. At 100
◦C
, the experiment continued for up to 5
h
to simulate a
longer retention time inside a continuous distillation process. Afterwards, the setup cooled down,
the round-bottom fasks were weighted, and the distillate and bottom products were withdrawn.
24
3 Methods
3.2 Experimental demonstration of the main COMET process units
3.2.1 Chemicals and materials
The reactants OME
1
(purity
>
99
.
8 %) and MeOH (purity
>
99
.
8 %) were purchased from
Brenntag GmbH and provided from ChemCom Industries B.V.. The reactant FA(aq.) was
provided as a stabilized aqueous FA (approx. 37
wt
%) solution with low amounts of MeOH
(
≤
0
.
5
wt
%) by ChemCom Industries B.V.. The catalyst A46 was provided by INAQUA
Vertriebsgesellschaft mbH. IER III was purchased from Merck Chemicals GmbH.
3.2.2 Analytics
The samples were analyzed by ASG Analytik-Service AG. The FA content was determined by the
sodium sulfte method for concentrations higher than 0
.
2
wt
% and by voltametric analysis for
smaller concentrations [124]. The H
2
O content was determined by Karl-Fischer titration. The
content of OME
1-10
, MeOH, TRI, tetroxane and MEFO was determined by a gas chromatographic
method using fame ionization detection (GC-FID). An online GC equipped with a thermal
conductivity detector (GC-TCD) was used for online measurements. The applied methods for
the GC analysis were ASG 2506 GC-FID for the organic compounds and ASG 2504 GC-FID for
TRI and tetroxane.
3.2.3 Concentrated FA(aq.) feed preparation
A cascade of two thin flm evaporators was used to provide the concentrated FA(aq.) solution.
The setup was provided and operated by VTA Verfahrenstechnische Anlagen GmbH & Co. KG
and directly connected to the OME synthesis setup. As a preparation, the stabilized FA(aq.)
containing about 37
wt
% FA was concentrated to about 55
wt
% FA, which is the concentration of
the product of a FA(aq.) production plant. The product was stored in a heated tank and further
concentrated in a frst step to about 75
wt
% FA and in a second step to about 85
�
89
wt
% FA.
The evaporators were operated under vacuum at about 200
�
600
mbar
and temperatures of the
heating fuid of 100 � 150 ◦C. A simplifed process fow diagram is illustrated in Figure 3.2.
Figure 3.2:
Simplifed process fow diagram for the concentration of an FA(aq.) to 85
�
88
wt
%
FA using a cascade of two evaporators operated at 200
�
600
mbar
and 100
�
150
◦C
heating fuid [4]. E, evaporator; P, pump.
25
3 Methods
3.2.4 OME synthesis
A setup for the synthesis of OME in a fxed bed reactor with a capacity of 1
�
5
Lh
�1
was used.
The setup contains a pump and heat exchanger to pressurise and heat up the reactant OME
1
to
meet the synthesis conditions, as well as a pump and heated tubes for the concentrated FA(aq.)
from the thin flm evaporators. After mixing the reactants, the stream was directly converted
in a fxed bed reactor flled with the catalyst A46 and heated to about 90
◦C
. The synthesis
product was mixed with additional MeOH, cooled to ambient temperature, depressurized, and
directly analyzed using the GC-TCD before it was stored at ambient temperature. A simplifed
process fow diagram is illustrated in Figure 3.3.
Figure 3.3:
Simplifed process fow diagram for the OME synthesis of OME
1
and concentrated
FA(aq.) solution over A46 for a capacity of 1
�
5
Lh
�1
at about 90
◦C
[4]. H, heat
exchanger; P, pump; R, reactor; T, tank; V, valve.
Before using the catalyst A46, it was stored in a mixture of FA, H
2
O, MeOH and OME
1
to
prevent further swelling inside the reactor unit.
The addition of MeOH to the OME synthesis product allowed for a stable storage and transport.
The stabilization was a preventive measure to ensure a homogeneous liquid solution without
precipitation even at low temperatures and long storage periods. The amount of additional
MeOH was determined to meet the demand for the reactive distillation column. This enabled
a high dilution without exceeding the demand of MeOH, which otherwise would change the
performance of the reactive distillation column. An illustration is presented in Figure 4.7. The
target of the reactive distillation column was a bottom product with a concentration of about
60
wt
% FA and H
2
O, and an almost complete conversion of MeOH, assuming that OME
2-3
are
converted to FA and OME
1
as described by eqn. 2.7, MeOH and FA are converted to OME
1
and H
2
O as described by eqn. 2.5, and the distillate product is the azeotropic mixture of OME
1
and MeOH at ambient pressure.
The OME reactor product was neutralized using IER III and directly separated in a distillation
column, as described in detail in the following sections.
3.2.5 OME synthesis product neutralization
The OME synthesis product was pumped through a neutralization bed of IER III at ambient
temperature and stored prior to the thermal separation in the distillation column.
26
3 Methods
3.2.6 Thermal separation in CO-1
A DN50 glass distillation column with one upper and one lower section of 70
cm
height each
was used for all thermal separation steps [125]. For the investigation of the frst separation step
in CO-1, the two sections of the column below and above the feed were flled with Montz 750
structured packings. A horizontal reboiler was used which prevented the fooding of the column
in the start-up phase due to strong foaming of the mixture at boiling conditions. The column
was continuously operated at ambient pressure with a feed rate of 1
�
2
Lh
�1
for about 250
kg
OME synthesis product for about 200
h
. After achieving steady state, distillate and bottom
products were withdrawn continuously from the column. The target of the separation was a
split between OME
2
and OME
3
and the removal of FA and H
2
O from the bottom product. A
simplifed process fow diagram is illustrated in Figure 3.4.
Figure 3.4:
Simplifed process fow diagram of the DN 50 glass distillation setup for a feed rate of
1
�
5
Lh
�1
[4]. CO, distillation column; H, heat exchanger; P, pump; VP, vacuum
pump.
3.2.7 Reactive distillation in CO-2
The core step of the COMET process concept takes place in CO-2. The distillate product of
CO-1 was used as the feedstock for CO-2. The same distillation setup of CO-1 was used for the
experimental investigation of the reactive distillation CO-2. The lower section of the column
was flled with a fxed bed of the catalyst A46 and Montz 750 structured packing on top. The
upper section above the feed was flled similarly with Montz 750 structured packing. The column
was continuously operated at ambient pressure with a feed rate of 0
.
5
�
1
Lh
�1
and for about
150
kg
feedstock under continuous withdrawal of distillate and bottom product at steady state
conditions. The target of the separation was a split between the azeotropic mixture of OME
1
and MeOH, and a mixture of FA and H2O, accompanied by the conversion of FA and MeOH.
3.2.8 Thermal separation in CO-3
The bottom product of CO-1 containing mainly OME
≥3
was fractionated in CO-3. The same
distillation setup of CO-1 was used for the experimental investigation of the thermal separation
27
3 Methods
in CO-3 with Montz 750 packings in the upper and lower section. The setup was continuously
operated at about 200
mbar
with a feed rate of 2
�
3
L/h
and for about 50
kg
feedstock under
continuous withdrawal of distillate and bottom product at steady-state conditions. The target of
the separation was a split between OME5 and OME6.
28
3 Methods
3.3 Process simulation and evaluation
This section addresses the general assumptions and system boundaries, the methodology of
process modelling and simulation, as well as process evaluation and comparison criteria.
3.3.1 General assumptions and system boundaries
The system boundaries were set for the evaluation on the simulation level from the feedstock H
2
and CO
2
, followed by the synthesis and purifcation of intermediates up to the desired product
OME
3-5
. It is assumed that the production plants are integrated in a chemical park where the
necessary infrastructure for the provision of utilities such as steam and cooling water is available.
CO
2
and H
2
are supplied from renewable non-fossil sources. The waste streams of the processes
are wastewater and exhaust gases. The production of OME3-5 consists of several sub-processes,
which are altogether material and heat integrated. The obtained distribution of OME chain
lengths in the fnal product OME
3-5
difers slightly between the investigated process routes.
Nevertheless, it is assumed that in all cases, the specifcation range is fulflled without further
processing and that the heating value from Held et al. [109] of 18
.
9
MJ kg�1
represents the
actual heating value as a good approximation. The system boundaries are illustrated in Figure
3.5.
Figure 3.5:
System boundaries set on the simulation level of the OME
3-5
production processes
[1].
3.3.2 Process modelling and simulation
Steady-state simulations for all OME
3-5
production processes were implemented using Aspen
Plus
®
sofware V11 and V12 from Aspen Technology Inc. Aspen Energy Analyzer V11 and V12
and Aspen Process Economic Analyzer V11 and V12 were used for heat integration and unit
operation dimensioning.
The components considered in the simulations were H
2
, CO
2
, CO, N
2
,O
2
, FA, MeOH, H
2
O,
OME1-10, HF1-10 and MG1-10. Most of these components such as the acetals, glycols and OME
29
3 Methods
are not included in the standard Aspen database, and thus new components were added in
the property analysis environment. An overview of the pure component thermodynamic and
thermophysical properties used in the simulations can be found in the Appendix in section A.3.
For the side products HF
n
and MG
n
in mixtures containing FA, H
2
O and MeOH, the true
composition was used for the process simulations, which considers the presence of HF
n
and MG
n
.
The overall composition, considering the stoichiometric decomposition of HF
n
and MG
n
to their
reactants MeOH, H
2
O and FA, was used for the evaluation and the presentation of the results.
OME, HF and MG are not included in the standard Aspen database. The thermodynamic
properties of these intermediates are described in the original literature but the right property
implementation and consideration in a process simulation require fundamental knowhow about
this reactive system. The work by Maurer, Hasse, Burger and Schmitz et al. [18, 80, 91, 92, 126]
ofers a concrete basis for the validation of the simulation models. Furthermore, the methodology
of the implementation of these reaction and phase behavior considerations along with the whole
fowsheet simulation was introduced by Bongartz et al. [108] employing tool boxes from the
electrochemical feld, namely, the chemistry section in Aspen Plus
®
and is further modifed in
this work. There have been several signifcant modelling and simulation eforts to describe this
complex system behavior, where the works by Burger et al. [127], Schmitz et al. [18, 80, 91,
92], Bongartz et al. [108], and Ouda et al. [128] are acknowledged, and the simulation results
generated in this work are a progression on their previous work.
To simulate purifcation processes, in particular thermal separations, interaction parameters are
required, which describe the real behavior of the gas and the liquid phases. For mixtures that
contain FA, a UNIFAC-based model was introduced by Maurer et al. [126]. This model was
extended in the following decades by adding new components and by adapting the interaction
parameters to new experimental data. Schmitz et al. [91] published a new version of the model
considering OME
n
which was adapted and implemented. An overview of the model and its
validation is presented in the Appendix in section A.3.2 and A.3.3. Adequate model parameters
and implementation of the thermodynamic model are crucial for a realistic simulation of this
special reactive mixture. A variety of reaction models describing the MeOH, FA(aq.), monomeric
FA, OME
1
and OME
n
syntheses were used to assess the product compositions exiting the reactors.
The models implemented in the simulation environment can be found in the process description
of the sub-processes. The synthesis of FA was described based on the conversion and yields from
literature values, while the other syntheses were described according to published kinetic models,
as discussed in the respective sections.
Initially, the sub-processes were implemented separately in the simulation platform. Aferwards,
the material integration interconnecting these sub-processes to describe the OME
3-5
production
processes was implemented. The production capacity was adjusted to 100
kt a�1
OME
3-5
, and
the heat integration was conducted to improve the overall process energy efciency. The heat
integration was conducted following the pinch method using Aspen Energy Analyzer. The
reactors for the MeOH and FA synthesis and the combustion, as well as the heat exchangers for
heating or quenching of the FA feed or product stream before and after the reactor were not
considered for the direct heat integration. In those heat exchangers the heat transfer coefcient
is signifcant for a fast heat exchange, therefore, steam is used as the utility. This steam was
30
3 Methods
supplied from or utilized in the process for further integration. After the design of the heat
exchanger network the fow sheet of the simulation was adjusted and step-by-step converged.
Since the total process is a combination of several sub-processes, more recycle loops that are
interconnected were converged. Standard numerical solvers in Aspen Plus
®
were used, while the
complex loop convergence of the total process was achieved stepwise starting with connecting
the sub-processes.
For the simulation of the reactive distillation column of the COMET process, the kinetic model
from Schmitz et al. [80] for the OME synthesis over A46 for feedstocks comprising MeOH, FA,
H
2
O and OME
1
was used. The model was implemented in a Fortran subroutine and activated on
the catalyzed trays inside the Aspen Plus process unit RadFrac column. The implementation of
the kinetic model was validated with the experimental results from Drunsel [129], who investigated
a reactive distillation column in the OME
1
production process for a similar reactive separation
task, however, using the catayst A15 instead of A46. A good agreement was obtained between
experimental and simulation results. Furthermore, the subroutine was slightly adjusted to be
used in a reactor unit and was validated with the experimental results from Schmitz et al. [80]
with a good agreement. In contrast to the kinetics of the OME formation as described by eqn.
2.5-2.7, the model assumes the formation of HF
n
and MG
n
as described by eqn. 2.1-2.4 to be in
chemical equilibrium at all retention times.
The formation of the side products TRI, DME, MEFO, FOAC and tetroxane was not considered
in the process simulations, due to very small concentrations in the synthesis product when A46
is used as a catalyst.
Pressure values presented in this work describe the absolute pressure.
3.3.3 Process evaluation and comparison criteria
The implemented OME
3-5
production processes were evaluated using various key performance
indicators (KPIs). These KPIs were translated into process evaluation criteria used to compare
the process routes, given that they provide a concise summary of the diferent process routes
in terms of mass and energy balance. The process route utility demands and process energy
efciencies were evaluated based on the total mass and energy balance.
In this work, the energy efciency of the overall process,
ηenergy
, is defned by eqn. 3.6, where
m
denotes the mass fow rate of the reactants
i
and the OME
3-5
containing product stream. LHV
is the lower heating value at 298
K
, while
Qk
and
Wl
represent the externally supplied heat
streams and electric power demand, respectively.
mOME3�5 · LHVOME3�5
ηenergy = (3.6)
k Qk + l Wl + i mi · LHVi
In addition to the energy efciency of the process routes, the material balance was assessed,
and the performance was indicated by two parameters. The frst parameter,
ηC
, refecting the
carbon efciency, i.e., the ratio of carbon atoms,
C
, in the feedstock and the carbon atoms in the
OME
3-5
product stream, is defned by eqn. 3.7. The second parameter,
ηmass
, considers the mass
31
3 Methods
fow rates, i.e., the ratio of the OME
3-5
product mass fow rate with respect to the feedstock
mass fow rate, as defned by eqn. 3.8. m denotes the mass fow rate of the reactants i.
COME3�5
ηC = (3.7)
i Ci
mOME3�5
ηmass = (3.8)
i mi
3.4 Process evaluation for an improved energy efciency of Power-to-X
processes using heat pumps
The evaluation of the overall energy efciency for the production of OME
3-5
based on extended
system boundaries follows the results for the process simulation of P1, presented in sections 5.3.2,
5.3.3 and 5.3.4, comprising the mass balance, energy demand and process efciencies, respectively.
Furthermore, parameters from public literature, and technical data sheets were used for the
consideration of the upstream processes.
The extended system boundary is illustrated in Figure 3.6, covering the OME
3-5
production
from all the primary resource conversion processes from electricity, seawater, and air to the
fnal OME product. The efciency of each sub-process considering mature technologies will be
discussed in section 5.4. OME
3-5
are produced from H
2
and CO
2
via the intermediate production
of MeOH and FA(aq.), the H
2
O electrolysis using renewable energies, as well as the upstream
seawater preparation, CO
2
preparation via direct air capture (DAC), heat pumps for upgrading
low-temperature exess streams and suitable applications for low-temperature heat. A description
of the extended system boundary discussing the assumptions and results for the individual
scenarios I-VI is presented in section 5.4.
Besides the system boundary, the defnition of the overall energy efciency of eqn. 3.6 was
extended for the the whole value chain to additionally consider excess heat streams
Qj
as
by-products of the OME3-5 production process, as described by eqn. 3.9.
mn · LHVn + j Qj
n
ηprocess = (3.9)
k Qk + l Wl + i mi · LHVi
32
3 Methods
Figure 3.6:
System boundary of a sustainable OME
3-5
production process from H
2
and CO
2
for
the evaluation of the overall energy efciency. In scenario I, the production of OME
3-5
is considered from H
2
and CO
2
via the intermediate production of MeOH and FA,
including heat integration. Scenarios II and III additionally cover the preparation
and provision of H
2
via H
2
O electrolysis and CO
2
via DAC. In scenario IV, HTHP
are included to lift low-temperature excess heat streams to usable temperature
levels. Scenario V additionally considers seawater desalination and scenario VI the
application of excess heat streams for other applications [2].
33
4 OME process description
4.1 OME3-5 production processes
Due to various potential suppliers of methyl groups and oxymethylene groups, diferent feedstock
combinations can be used to produce OME
3-5
. Depending on the formation of H
2
O as a side
product during the synthesis of OME
n
, the reaction systems are classifed as anhydrous or
aqueous. Methyl group suppliers for anhydrous reaction systems are mainly DME and OME
1
,
because their conversion to OME
n
only requires a chain propagation with oxymethylene groups,
as described by eqn. 2.7 and presented in Table 4.1. Oxymethylene group suppliers for anhydrous
reaction systems are mainly TRI and monomeric FA, due to the absence of H
2
O or MeOH.
Formalin, concentrated FA(aq.) or pFA contain H
2
O and are, therefore, used as oxymethylene
group suppliers for aqueous reaction systems.
Table 4.1:
Methyl group suppliers and oxymethylene group suppliers for anhydrous and aqueous
OME reaction systems [4].
methyl group supplier oxymethylene group supplier
anhydrous DME, OME1 TRI, monomeric FA
aqueous MeOH, DME, OME1 FA(aq.), pFA, TRI, monomeric FA
Feedstocks containing MeOH generally lead to the formation of H
2
O as a side product in the
aqueous OME synthesis, as described by eqn. 2.5 and 2.6. This H
2
O needs to be separated
and extracted from the process loop to prevent accumulation. Figure 4.1 shows a simplifed
scheme for the production of OME
3-5
from various feedstocks. It contains a reactor for the
OME synthesis R, two distillation columns CO-1 and CO-2 for product purifcation and a H
2
O
separation unit S for aqueous reaction systems. For the H
2
O separation various methods were
proposed, such as extraction, adsorption or membrane, as discussed in section 2.5.
Various process concepts for the production of OME
3-5
were proposed in patents and other
publications and some of them are realized in large-scale production plants in China. However,
information are scarce about their performance, the quality and composition of the fnal OME
product and the long-term operation [89, 106]. The main OME
3-5
production processes discussed
in the literature are summarized in Table 4.2, focusing on the feedstock, main advantages, and
hurdles. A detailed description is provided in the following sections. A comparison with the
COMET process based on their performances in terms of OME
3-5
yield, energy demand and
technical feasibility is discussed in section 5.3.5. Further process concepts were proposed in the
literature which show signifcant disadvantages in comparison to the process concepts presented
in Table 4.2, as discussed in section 4.1.9.
35
4 OME process description
Figure 4.1:
OME
3-5
production process for various feedstocks, following aqueous and anhydrous
reaction systems [4]. CO, distillation column; R, reactor; S, H2O separator.
4.1.1 OME1 and TRI (anhydrous synthesis)
Schelling et al. [13] proposed a process concept for the anhydrous synthesis and purifcation of
OME
3-5
from OME
1
and TRI, which was updated by Burger et al. [14, 15] reducing the number
of distillation columns to two. The feed mixture of OME
1
and TRI is mixed with the recycle
streams and synthesized in the reactor as illustrated in Figure 4.1. The OME synthesis proceeds
fast, as shown in Figure 4.2a, which presents the experimental result of the OME synthesis from
OME
1
and TRI by Burger [14]. Furthermore, a comparatively high selectivity for OME
3-5
is
obtained with very low side product formations using the catalyst A46, as shown in Figure 4.2b.
The choice of catalyst is crucial for the reaction kinetics and side product formation. Using
A36, Burger detected the side products MEFO and DME, while only traces of MEFO could be
detected with A46 at temperatures higher than 75
◦C
. The investigations in this work regarding
the OME synthesis from OME
1
and TRI over various catalysts also show comparatively high
side product formations for zeolites but low side product formations for A46, Dowex and Nafon
at 60 ◦C, as discussed in section 5.1.3.
After the reactor, the synthesis product mainly containing OME
1-10
, FA and TRI is separated in
a frst distillation column to separate OME
≥3
from the more volatile components OME
1-2
, FA
and TRI. The distillate product is recycled to the reactor and the bottom product is separated in
a second distillation column to provide the target product containing OME
3-5
which is separated
from the process and a bottom product containing OME≥6 which are recycled to the reactor.
The advantages of this process concept are the simple design and no formation and, therefore,
necessary separation of H
2
O from the loop. However, Lautenschütz [25] investigated in a blank
experiment the conversion of OME
1
alone in presence of the catalyst A36. Besides OME
1
, the
product mixture contained 2
wt
% MeOH and 3
wt
% OME
2
. In a subsequent experiment he
dried OME
1
before the addition of A36 and no MeOH or OME
2
could be detected in the product
mixture. This emphasizes the need for a very dry feedstock to prevent the formation and fnally
accumulation of H2O inside the process loop.
36
4 OME process description
Table 4.2: Advantages and main hurdles of various OME3-5 production process concepts [4].
Anhydrous synthesis Aqueous synthesis
Feedstock
OME
1
and DME and OME
1
and MeOH and MeOH and OME
1
and
TRI [13–
15]
TRI [16,
17]
monomeric
FA [1]
FA(aq.)
[18, 19]
monomeric
FA [1]
FA(aq.) or
pFA [1]
(+)
main
advan-
tages
and
(-) main
hurdles
+ high
OME
3-5
yield after
the synthe-
sis
+ simple
product
purifcation
- complex
and energy-
intensive
preparation
of TRI [87,
94, 109,
130]
+ DME
is cheaper
than OME
1
[119]
- complex
and energy-
intensive
preparation
of TRI [87,
94, 109,
130]
- high
MEFO
selectiv-
ity [84, 88,
131–133]
+ high
OME
3-5
yield after
the synthe-
sis
+ poten-
tially sim-
pler and
cheaper
produc-
tion of
monomeric
FA
- very low
TRL of the
monomeric
FA produc-
tion
+ com-
paratively
cheap feed-
stock
- formation
of H
2
O
as a side
product
- low
OME
3-5
yield after
the synthe-
sis
- low TRL
of the H
2
O
separation
methods
+ com-
paratively
cheap feed-
stock
- similar
hurdles to
MeOH and
FA(aq.)
- very low
TRL of the
monomeric
FA produc-
tion
+ fairly
high
OME
3-5
yield after
the synthe-
sis
- similar
hurdles to
MeOH and
FA(aq.)
The main disadvantage of this process concept is the preparation of the feedstock TRI which is
complex and energy-intensive, mainly due to the low conversion of FA to TRI in the reactor and
the resulting high recycle streams [87, 94, 109, 130].
In comparison to alternative process routes the conversion of OME
1
and TRI to OME
3-5
shows
the advantage of a high OME
3-5
yield which leads to an increase of the mass fraction of OME
3-5
from 5 to 34
wt
% before and after the reactor [109]. Therefore, comparatively small recycle rates
are obtained in the loop and the two distillation columns for the product purifcation require
a low heat demand of about 8 % in comparison to the energy content of the OME
3-5
product
based on the LHV, as presented in Table 5.5 [109]. However, this advantage is outweighed by
the energy-intensive feedstock preparation, resulting in an overall energy efciency of 29
�
37 %
considering the entire process chain from H
2
O electrolysis and CO
2
capture via the MeOH, FA,
TRI and OME1 production towards the fnal OME3-5 product mixture [109].
Considering the possibility of scale-up for a sustainable OME
3-5
production based on OME
1
and TRI, the sub-processes for the production of MeOH (from CO
2
and H
2
), FA, TRI and
OME
1
are state-of-the-art and, therefore, show high technology readiness level (TRL) [1, 134,
135]. Furthermore, the sub-process of the OME
3-5
production only consists of state-of-the-art
process units. Therefore, this process is an energy-intensive but feasible process concept for the
production of OME3-5 in the near future.
37
4 OME process description
(a) (b)
Figure 4.2:
OME synthesis from OME
1
and TRI over A46 (conditions: OME
1
/TRI = 2
.
85
g g�1
,
A46/(OME
1
+TRI) = 0
.
8
wt
%, 65
◦C
, batch) by Burger [4, 14]. (a) shows the
reaction progress and (b) shows the equilibrium composition. The values describe
the mass fractions of the synthesis products.
4.1.2 DME and TRI (anhydrous synthesis)
Ströfer et al. [16, 17] proposed a process concept for the anhydrous synthesis and purifcation of
OME
3-5
from DME and TRI similar to the process concept for OME
1
and TRI, as illustrated in
Figure 4.1. Due to the high vapor pressure of DME, the synthesis and frst distillation column
are operated at elevated pressure levels, which leads to a far higher reboiler temperature of about
300
◦C
in comparison to alternative processes with temperatures around 200
◦C.
This results
in more expensive heat sources. Furthermore, the stability regarding thermal decomposition of
OME should be experimentally investigated at this high temperature level. The main advantage
of this process concept is the absence of H
2
O. Furthermore, DME is a cheaper feedstock than
OME
1
[119]. However, the main disadvantage is the complex and energy-intensive preparation of
TRI [87, 94, 109, 130].
Considering the experimental results of Haltenort et al. [84, 88], Drexler et al. [131, 133] and
Breitkreuz et al. [132] regarding the OME synthesis from DME and TRI, the synthesis proceeds
slow, as illustrated in Figure 4.3a. Furthermore, a lower yield of OME
3-5
is obtained with very
high side product formations for various catalyst systems and already at comparatively low
temperatures of 80
◦C
, as illustrated in Figure 4.3b [131, 132]. The highest OME
3-5
concentration
was obtained after 76
h
, whereas the mixture still contained a high share of unreacted feedstock.
For longer retention times, the concentration of OME
3-5
reduced, due to an increasing formation
of MEFO and FOAC. Unfortunately, the catalyst A46, which shows very small side product
formations in the OME synthesis from OME
1
and TRI, is not active for this feedstock, due to
its low acid concentration [132].
Besides OME, especially MEFO is produced with an increasing yield. For long retention times
the yield of MEFO exceeds the yield of the product OME
3-5
. Therefore, the process concept
needs to be adjusted to include the separation and purifcation of MEFO as a valuable side
product. However, due to the narrow boiling point between OME
1
and MEFO a high purity
separation using distillation columns is challenging [136, 137]. This also complicates a sustainable
38
4 OME process description
(a) (b)
Figure 4.3:
OME synthesis from DME and TRI over A36 (conditions: DME/TRI = 1
.
80
gg
�1
,
A36/(DME+TRI) = 4
.
4
wt
%, 90
◦C
, batch) by Breitkreuz [4, 132]. (a) shows
the reaction progress and (b) shows the composition after 76
h
with the highest
concentration of OME
3-5
. The values describe the mass fractions of the synthesis
products.
production of OME3-5 based on DME and TRI in the near future.
4.1.3 DME and monomeric FA (anhydrous synthesis)
To date there is no process concept published for the OME production from DME and monomeric
FA. However, this process was investigated in the frame of the NAMOSYN project in which
diferent OME
3-5
production processes were evaluated and compared [1, 138]. The process
concept for the production of OME
3-5
from DME and monomeric FA is similar to the process
from OME
1
and TRI, as illustrated in Figure 4.1. Furthermore, the same advantage of an
anhydrous synthesis can be obtained, which potentially leads to comparatively high OME
3-5
yields. The feedstock DME is cheaper than OME1 and the production of monomeric FA shows
potential to become simpler and cheaper than the production of TRI. In comparison to the
partial oxidation of MeOH, as described in eqn. 4.4, the dehydrogenation of MeOH towards FA
produces H
2
instead of H
2
O as a by-product [110]. H
2
can be separated and recycled to the
MeOH synthesis, which results in a stochiometrically lower H
2
demand and potentially reduces
the production costs, since the feedstock H
2
generally has the biggest share on the production
costs for various PtX products, considering a sustainable production [1, 119]. Moreover, using
monomeric FA instead of TRI might reduce the formation of side products in comparison to the
OME synthesis based on DME and TRI. However, the main disadvantages are a very low TRL
of the anhydrous FA synthesis and many open investigations regarding its usage for the synthesis
of OME [1]. So far, the anhydrous FA synthesis was investigated in laboratory experiments
[110, 139]. Investigations regarding catalyst deactivation and long-term experimeriments are
still to be successfully completed before its demonstration. Furthermore, besides the synthesis,
the monomeric FA product is gaseous and needs to be absorbed from the synthesis product
mixture without using H
2
O or MeOH as a washing liquid, which are used for the FA(aq.)
39
4 OME process description
separation but would lead to the formation of many side products in the OME synthesis, as
described by eqn. 2.1-2.4. The washing liquid should either be DME, the recycle stream con-
taining the volatile components in the OME
3-5
sub-process, or a separate component which
does not react in the OME synthesis and can be separated and recycled to the absorption
column. The solubility of monomeric FA in DME, OME or other suitable candidates should
be investigated at suitable conditions for the absorption, which strongly difers between the
washing liquids, to provide a liquid product stream. Therefore, a demonstration and scale-up for a
sustainable production of OME
3-5
based on DME and monomeric FA is unlikely in the near future.
4.1.4 OME1 and monomeric FA (anhydrous synthesis)
Within the scope of this work, a process concept is proposed, simulated and evaluated for the
anhydrous synthesis and purifcation of OME
3-5
from OME
1
and monomeric FA, similar to the
process concept for OME
1
and TRI as illustrated in Figure 4.1 [1]. The process concept is
described in this section, whereas the results and evaluation is discussed in section 5.3.5.
The process concept includs a H
2
O separation unit for the distillate product of the frst distillation
column to separate traces of H
2
O, which entered the process from the OME
1
feedstock. This
unit can also be omitted if a high purity OME
1
feedstock can be provided. In comparison to the
feedstock DME, OME
1
is more expensive [119]. However, synthesis experiments with similar
feedstocks of OME
1
and TRI as well as OME
1
and pFA show comparatively high selectivities
and low side product formations, if the temperature is kept below 80
◦C
and a suitable catalyst
system is used [14, 140]. Peter et al. [141] investigated the synthesis of OME
1
and gaseous
monomeric FA and found a comparatively high selectivity towards OME
3-5
with low side product
formations of MEFO and TRI, see Figure 4.4. With a smaller ratio of OME
1
to FA the yield of
OME3-5 can be further increased.
Figure 4.4:
OME synthesis from OME
1
and monomeric FA over OMe
3+
BF
4-
in EMIM
+
BF
4-
(ionic liquid) (conditions: OME
1
/FA = 1
.
58
g g�1
, OMe
3+
BF
4-
/OME
1
= 2
�
3
mol
%,
45
◦C
, continuous addition of gaseous FA) by Peter et al. [4, 141]. The values describe
the mass fractions of the synthesis product.
Besides the advantage of a potentially simpler and cheaper production of monomeric FA in
comparison to TRI, similar challenges to the OME production based on DME and monomeric FA
40
4 OME process description
need to be investigated. The provision of the feedstock in the liquid phase and the demonstration
of the anhydrous FA synthesis. The process simulation in this work assumes that the recycle
stream of the volatile components from the OME
3-5
sub-process is used to absorb monomeric FA
from the FA synthesis mixture. However, it is pointed out that the absorption of monomeric
FA is a crucial process step, and the assumption of a good solubility should be experimentally
investigated. Peter et al. [141] observed that the addition of gaseous FA to OME
1
without the
presence of a catalyst led to instant polymerization. This indicates a low solubility of monomeric
FA in OME
1
. Zimao et al. [142] on the other hand emphasized a good solubility of FA in OME
2
.
Regarding the process performance, the conversion of OME
1
and monomeric FA shows the
potential of a high OME
3-5
yield which leads to an increase of the mass fraction of OME
3-5
from
5 to 29
wt
% before and after the reactor, as presented in section 5.3.5. This leads to small recycle
rates and results in a heat demand for the two distillation columns of about 11 % in comparison
to the energy content of the fnal OME
3-5
product based on the LHV. Considering the assumption
from Held et al. [109] regarding H
2
O electrolysis and CO
2
capture, results in an overall energy
efciency of 27
�
36 %, including the production of the intermediate products MeOH, FA and
OME
1
. However, due to the low TRL of the monomeric FA production a demonstration and
scale-up of the sustainable production of OME
3-5
based on OME
1
and monomeric FA is unlikely
for the near future.
4.1.5 MeOH and FA(aq.) (aqueous synthesis)
Schmitz et al. [18, 19] proposed a process concept for the aqueous synthesis and purifcation of
OME
3-5
from MeOH and concentrated FA(aq.). The process concept is similar to the concept
for OME
1
and TRI, with the addition of a H
2
O separation unit, as illustrated in Figure 4.1.
The feed mixture of MeOH and concentrated FA(aq.) is mixed with the recycle streams and
synthesized in the reactor. The OME synthesis proceeds fast, as shown in Figure 4.5a, which
illustrates the experimental result of the OME synthesis from MeOH and pFA by Schmitz et al.
[80]. However, due to the presence of H
2
O and MeOH, FA reacts to HF and MG, as described
by eqn. 2.1-2.4. Therefore, a comparatively low selectivity of OME
3-5
is obtained, as shown in
Figure 4.5b. This can be increased by adding more FA to MeOH, but the fraction of OME
3-5
stays signifcantly smaller in comparison to the anhydrous routes. Using the catalyst A46, only
traces of MEFO and TRI were detected [27].
After the reactor, the synthesis product mainly contains FA, H
2
O, MeOH, HF, MG and OME
1-10
and is separated in a frst distillation column to separate OME
≥3
from the more volatile
components FA, H
2
O, MeOH, HF, MG and OME
1-2
. The distillate product is sent to a H
2
O
separation unit and afterwards recycled to the reactor. The bottom product is separated into the
target product OME
3-5
, which is separated from the process, and OME
≥6
, which are recycled back
to the reactor. The main advantage of this process concept is the feedstock, whose preparation
is simpler than the feedstock for the anhydrous routes. However, the main disadvantage is the
formation of H
2
O as a side product, which decreases the selectivity towards OME
3-5
and needs
41
4 OME process description
(a) (b)
Figure 4.5:
OME synthesis from MeOH and pFA over A46 (conditions: FA/MeOH = 0
.
89
gg
�1
,
A46/(MeOH+pFA) = 1
.
9
wt
%, 60
◦C
, batch) by Schmitz et al. [4, 80]. (a) shows
the reaction progress and (b) shows the equilibrium composition. The values describe
the mass fractions of the synthesis products.
to be separated from the loop. Schmitz et al. [18] proposed the utilization of adsorbents or a
membrane to separate and extract the side product H2O, which is described in section 2.5.
In comparion to alternative process concepts, the conversion of MeOH and FA(aq.) to OME3-5
shows a low OME
3-5
yield which leads to an increase of the mass fraction of OME
3-5
from 0
to 15
wt
% before and after the reactor [109]. Therefore, comparatively large recycle rates are
obtained in the loop. The purifcation of the synthesis product is energy-intensive in the two
distillation columns, whose reboiler duties sum up to about 47 % of the energy content of the
OME
3-5
product based on the LHV [109]. However, these disadvantages are outweighed by the
comparatively simple feedstock preparation, resulting in an overall energy efciency of 25
�
31 %,
considering the entire process chain from H
2
O electrolysis and CO
2
capture, via the MeOH and
FA production, towards the fnal OME
3-5
product mixture, as presented in Table 5.5. The TRL
of the production of the intermediate products MeOH and FA is very high and does not limit the
scale-up of a sustainable OME
3-5
production based on MeOH and FA. Recently a plant was built
to demonstrate the production of OME
3-5
from MeOH and FA covering all required process units
and enabling the separation of H
2
O using a membrane [122, 143]. The membrane is the main
bottleneck for a fast scale-up of this process concept, which, considering the application as a fuel,
will easily grow above 100
kt a�1
OME
3-5
for a single production plant, which results in about
24
kt a�1
H
2
O to be separated from the distillate stream of about 520
kt a�1
, see section 5.3.2.
4.1.6 MeOH and monomeric FA (aqueous synthesis)
Within the scope of this work, a process concept is proposed, simulated and evaluated for the
aqueous synthesis and purifcation of OME
3-5
from MeOH and monomeric FA, similar to the
process concept from MeOH and concentrated FA(aq.), as illustrated in Figure 4.1 [1]. The
process concept is described in this section, whereas the results and evaluation is discussed in
section 5.3.5.
42
4 OME process description
The main advantages of this process concept are a simple preparation for the feedstock MeOH
and a potentially simple preparation of the feedstock monomeric FA. The main disadvantages are
the presence of H
2
O in the OME synthesis and the low TRL of the monomeric FA production.
In comparison to the OME production based on MeOH and FA(aq.), the OME
3-5
yield is
slightly improved which leads to an increase of the mass fraction of OME
3-5
from 3 to 19
wt
%
before and after the reactor, see Table 5.5. This decreases the recycle rates. However, the
purifcation of the synthesis product is still energy-intensive in the two distillation columns, with
a heat demand of about 48 % of the energy content of the OME
3-5
product based on the LHV.
Considering the entire process chain starting from H
2
O electrolysis and CO
2
via the production
of the indermediate products MeOH and FA towards the target product mixture OME
3-5
and
considering the assumptions from Held et al. [109] regarding the electricity and heat demand for
the H
2
O electrolysis and CO
2
preparation, an energy efciency of 28
�
37 % can be achieved.
Due to the low TRL of the monomeric FA production and the necessity of a H
2
O separation
unit, a fast demonstration and scale-up of a sustainable production of OME
3-5
based on MeOH
and monomeric FA is unlikely in the near future.
4.1.7 OME1 and FA(aq.) or pFA (aqueous synthesis)
Hackbarth et al. [89] published a list of OME production plants in China from which most of
them are based on the feedstock OME
1
and pFA. The process concept is similar to the OME
production process from OME
1
and TRI with the addition of a H
2
O separation unit, as illustrated
in Figure 4.1. Experimental results from Liu et al. [140] regarding the OME synthesis from
OME
1
and pFA show that a comparatively high yield towards OME
3-5
can be achieved, see
Figure 4.6. This can still be increased by increasing the ratio of OME
1
to pFA [140]. However,
they also reported that comparatively high temperatures of about 90
◦C
are benefcial for the
depolymerization of pFA, which lead to a high formation of side products, i.e. MEFO and DME,
using the catalyst NKC-9.
Figure 4.6:
OME synthesis from OME
1
and pFA over NKC-9 (conditions: OME
1
/pFA
= 4
.
31
g g�1
, NKC-9/(OME
1
+pFA) = 5
wt
%, 3h, 90
◦C
, batch) by Liu et
al. [4, 140]. The values describe the mass fractions of the synthesis product.
The depolymerization of the solid pFA can be accelerated using concentrated FA(aq.) instead,
which is an intermediate product for the pFA production and can be prepared using a cascade
43
4 OME process description
of evaporators, as discussed in the following section. The liquid concentrated FA(aq.) product
mainly consist of MG which also need to depolymerize as described by eqn. 2.4, but from a
smaller degree of polymerization. A disadvantage is the higher amount of H
2
O in the feedstock,
which reduces the selectivity towards OME3-5.
The main advantage of the process concept is the comparatively high selectivity of OME
3-5
,
which is increasing with decreasing H
2
O contents in the FA feedstock. Furthermore, production
processes for the feedstocks are state-of-the-art. The main disadvantage is the presence of H
2
O
in the OME synthesis, which needs to be separated from the loop.
In comparison to the OME production from MeOH and FA(aq.) the selectivity of OME
3-5
slightly
increases, resulting in an increase of the mass fraction of OME
3-5
from 4 to 19
wt
% before and
after the reactor, see Table 5.5. This leads to a heat demand of about 26 % for the reboiler of
the distillation columns in comparison to the energy content of the OME
3-5
product based on
the LHV. This is considerably lower than the heat demand for the OME production based on
MeOH and FA(aq.). Considering the entire process chain starting from H
2
O electrolysis and
CO
2
via the production of the indermediate products MeOH, FA and OME
1
towards the target
product mixture OME
3-5
and considering the assumptions from Held et al. [109] regarding the
electricity and heat demand for the H
2
O electrolysis and CO
2
preparation, an energy efciency
of 26
�
32 % can be achieved. Similar to the production of OME
3-5
from MeOH and FA(aq.),
the H2O separation unit is the bottleneck for the scale-up of this process concept.
4.1.8 COMET (aqueous synthesis)
The COMET process concept [144] was developed within the framework of this thesis. It is
based on the commercially available MeOH and FA(aq.) feedstock and produces mainly high
purity OME
3-5
, as illustrated in Figure 4.7. For the separation of H
2
O from the loop, a reactive
distillation column is used.
The COMET process starts at the concentration of FA(aq.) (stream 1), which can be the product
stream of a state-of-the-art FA production process with a concentration of 50
�
55
wt
% FA [1,
134]. Before the concentration, the stream is mixed with the distillate of the second evaporator
E-2 and the bottom of the third evaporator E-3. FA(aq.) is usually concentrated in a cascade of
two evaporator stages E-1 and E-2 to provide a concentrated FA(aq.) of 85
�
88
wt
% FA (stream
5) and an aqueous stream containing 10
�
25
wt
% FA (stream 3), depending on the feed mixture
and concentrated FA(aq.) product composition. The concentrated FA(aq.) (stream 5) is used for
the production of OME and mixed with the recycle streams, containing the azeotropic mixture
of OME
1
and MeOH (stream 10) and OME
≥ 6
(stream 14). The mixture is converted in a fxed
bed reactor R flled with an acidic heterogeneous catalyst. In contrast to the OME production
process based on MeOH and FA(aq.) [18], the reactor inlet stream contains mainly OME1 as a
methyl capping source. This improves the selectivity towards OME
3-5
. The comparatively high
selectivity further increases with decreasing H
2
O and MeOH concentrations in the concentrated
FA feedstock (stream 5) and OME
1
recycle (stream 10). The synthesis product mixture mainly
containing FA, H
2
O, MeOH and OME
1-10
(stream 7) is separated in a cascade of three distillation
44
4 OME process description
Figure 4.7:
COMET process concept for the production of OME
3-5
from MeOH and FA(aq.)
feedstocks [4]. The light grey arrows and process units were added in this work to the
FA concentration sub-process to improve the recycle of FA. CO, distillation column;
E, evaporator; R, reactor.
columns. In the frst distillation column CO-1, OME
≥ 3
are separated from the more volatile
components FA, H
2
O, MeOH and OME
1-2
. Thereby, OME
3
cannot be completely separated to
the bottom product, a small fraction remains in the distillate. In the third distillation column
CO-3, OME
≥ 6
are separated and recycled to the reactor to provide the fnal product (stream
13) as a mixture of OME
3-5
. The distillate product of CO-1 is mixed with MeOH (stream 9)
and sent to a reactive distillation column CO-2, to separate an azeotropic mixture of OME
1
and
MeOH (stream 10) from FA and H
2
O (stream 11). On the catalytic trays, two main conversions
take place. First, OME
2-3
are converted to OME
1
and FA over an acidic heterogeneous catalyst,
as described by eqn. 2.7. Besides, MeOH and FA are converted to OME
1
and H
2
O, following
the acetalization reaction as described by eqn. 2.5. The mechanism on the catalytic trays is
illustrated in Figure 4.8. Due to the evaporation and, therefore, the separation of the volatile
product OME
1
from the liquid reaction mixture, the equilibrium of eqn. 2.5 and 2.7 shifts, and
the reactions proceed towards the production of OME
1
. Therefore, with sufcient retention time,
OME
2
is converted to a large extent to OME
1
, while the conversion of FA towards OME
1
is
limited by the amount of MeOH. The mixture is separated into the azeotropic mixture of OME
1
and MeOH in the distillate (stream 10) and a mixture of FA and H
2
O in the bottom (stream
11). The distillate is recycled back to the reactor and the bottom product is recycled to the
evaporator E-2 for the FA concentration to separate H
2
O from the process and recycle FA back
towards the OME reactor. Therefore, the reactive distillation column prevents the accumulation
of H
2
O inside the loop and solves the challenging H
2
O management. In contrast to the H
2
O
separation from the loop using adsorption or membranes, in the COMET process H
2
O is not
separated selectively but together with the remaining FA. This signifcantly reduces the risk
of precipitation, since enough H
2
O is left to convert the remaining FA to comparatively short
45
4 OME process description
chain MG
n
which stay liquid at elevated temperature for sufcient retention time to downstream
processing steps.
Figure 4.8:
H
2
O separation from the COMET process via reactive distillation [4]. The left
side shows the reactive distillation column with the main components of the feed
and product streams. The illustration on the right side shows the interaction on a
catalytic tray and was adopted from Schmitz et al. [91].
A similar concept for a reactive distillation column is applied in the OME
1
production process
by Drunsel et al. [129, 145] with the purpose to achieve a complete conversion of FA after the
OME1 reactor.
The amount of MeOH (stream 9) added to the feed of the reactive distillation column CO-2
defnes the conversion of FA and oxymethylene groups with MeOH on the catalytic trays towards
OME
1
, following eqn. 2.1, 2.5 and 2.7. Therefore, a variation of the amount of MeOH (stream
9) varies the amount of OME
1
produced as the distillate product of the reactive distillation
column CO-2. For the OME synthesis in the fxed bed reactor R a constant ratio of OME
1
to
concentrated FA(aq.) before the reactor is required. Therefore, the amount of MeOH (stream 9)
can be defned to exactly produce the amount of OME
1
required for the OME synthesis. Or
the amount of MeOH (stream 9) can be increased to produce more OME
1
than required by the
OME synthesis and the excess OME
1
can be extracted as a by-product. Another advantage
of the COMET process is that the process ofers a tunable product portfolio of OME. In the
present work the amount of MeOH added to the reactive distillation column CO-2 was limited
to only produce the required amount of distillate product (stream 10) for the OME synthesis
and, therefore, achieve higher OME
3-5
selectivity. Considering the production of OME
1
as a
side product of the COMET process, another distillation column can be added to achieve high
purities of the OME
1
side product, similar to the second distillation column of the production
process for OME1 [129].
Besides the OME
3-5
product (stream 13), wastewater (stream 3) is produced with FA concen-
trations of about 10
�
25
wt
%. This by-product stream is not limited to the COMET process
but part of all OME
3-5
production processes using FA(aq.) as an intermediate product. Instead
of its disposal and to increase the carbon yield of the process, several strategies are possible to
handle this stream. In the present work, the stream was partly send to the absorber column
46
4 OME process description
of the FA(aq.) production and concentrated in an additional distillation column CO-4. The
concentrated FA stream was further concentrated in another evaporator E-3 to recycle the
concentrated FA stream and to separate the stream with a low FA concentration (stream 2).
This stream is also the purge stream for traces of MeOH and other volatile components to avoid
accumulation in the loop. Instead of its disposal, this stream can be used to dilute an FA(aq.)
product stream to prepare a stable formalin product.
The performance of the COMET process concept was simulated and evaluated and is presented
in section 5.3.5 together with a comparison to alternative OME3-5 production processes.
4.1.8.1 Expanding the system boundary to H2 and CO2 feedstocks including the intermediate
production of FA and MeOH
To enable a consistent basis of comparison with alternative OME
3-5
production processes, the
system boundary is extended to account for a sustainable OME
3-5
production based on green H
2
and captured CO
2
. The intermediate production of MeOH and FA(aq.) is described in detail in
section 4.2. A simplifed process fow diagram of the extended COMET process concept starting
from H2 and CO2 is illustrated in Figure 4.9.
Figure 4.9:
COMET process concept for the production of OME
3-5
from H
2
and CO
2
feedstock
with the intermediate production of MeOH and FA [4].
4.1.9 Further process concepts for the production of OME3-5
The processes described above for the production of OME
3-5
are comparatively simple and
efcient and contain the potential of a comparatively fast scale-up, after the main bottlenecks are
overcome and the feasibility is successfully demonstrated. However, various process alternatives
were published, which are more complex, contain unrealistic confgurations or redundant feedstock
combinations.
An OME
3-5
production in China is based on the feedstock MeOH and TRI, which has the
disadvantage of the energy-intensive TRI production and still requires a H
2
O separation unit
47
4 OME process description
inside the OME
3-5
sub-process [25, 89]. Therefore, the OME
3-5
production based on OME
1
and
TRI is simpler and it can already be scaled up.
Palkovits et al. [20, 21] proposed a process concept for the production of OME
3-5
based on MeOH
or OME
1
and FA. For the separation of OME
1-2
from H
2
O and MeOH, OME
1-2
are adsorbed on
activated carbon or hypercrosslinked polymers. However, H
2
O still needs to be separated from
the loop, preferably from MeOH and FA, to obtain higher yields of OME3-5.
Hagen et al. [22] proposed a process for the production of OME
3-5
based on DME and FA.
DME is used to produce FA and after the separation of DME, OME are formed and separated
in a reactive distillation column. In the described confguration it is unlikely that OME
≥2
are
formed and separated from the reactive distillation column in satisfying yields. Drunsel [129]
investigated a similar feed mixture in a reactive distillation column to produce OME
1
without
reporting the presence of OME≥2.
Qiang et al. [23] proposed a process for the production of OME
3-5
based on OME
1
and OME
≥6
.
The main advantages are an anhydrous synthesis without the need to separate H
2
O, high yields of
OME
3-5
and, therefore, a simple product purifcation. However, the availability of the feedstock
OME
≥6
is comparatively low since it is a by-product of the OME synthesis and there usually
recycled back to the reactor.
Furthermore, OME
3-5
production processes whose main bottleneck is the separation of H
2
O
from the loop can certainly already be constructed and scaled-up, if the operators accept and
handle large by-product streams which still contain signifcant amounts of unreacted feedstock
and OME
2
. This can be attractive if other processes are available which can use this by-product
stream as a feedstock, such as the process for the production of OME
1
. However, the scale
of the by-product streams would exceed the OME
3-5
product stream, which would result in
signifcantly lower yields of OME
3-5
. Considering the application of OME
3-5
in the mobil-
ity sector as a diesel fuel additive or alternative, many large-scale plants are needed, which
would very soon exceed the demand for the products of the by-product stream handling processes.
4.2 Upstream processes for the OME production based on H2 and CO2
4.2.1 MeOH production
MeOH synthesis is one of the oldest thermochemical processes with the highest production
capacities and is mainly based on fossil feedstocks. The development of a Cu-based process
enabled a signifcant reduction in the synthesis conditions to temperatures as low as 200
�
280
◦C
and pressure of 50
�
125
bar
. Advanced catalyst development allows MeOH synthesis based on a
CO
2
-rich feed with enhanced catalyst water tolerance [146, 147]. The process conditions for the
synthesis of MeOH are based on the work by Otto [148] and Bongartz et al. [107]. The reaction
network considered in this work is introduced in eqn. 4.1-4.3. The synthesis was simulated in an
isothermal plug-fow reactor considering the kinetics by Nestler et al. [149] and steam production
for cooling. The MeOH synthesis process takes place at 250
◦C
and 70
bar
in a fxed bed reactor
and the downstream purifcation of MeOH from non-reacted gases such as CO
2
, CO, H
2
and
48
4 OME process description
H
2
O goes through a cascade of fash drums with intermediate cooling, operating at diferent
pressure levels, then followed by a distillation column. The light gases with the non-reactants
are recycled back to the reactor to increase the product yield. A simplifed fowsheet of this
sub-process is shown in Figure 4.10.
CO2 +3H2 CH3OH + H2O (4.1)
CO + H2O CO2 + H2 (4.2)
CO +2H2 CH3OH (4.3)
Figure 4.10:
Simplifed process fow diagram for the production of MeOH from H
2
and CO
2
based on [4, 107].
4.2.2 FA(aq.) production and concentration
The FA(aq.) sub-process comprises the conversion of MeOH to FA. Formalin or pFA is synthesized
commercially from MeOH. The former is produced either via the silver catalyst based process or
the FORMOX process [134]. In these processes, MeOH is partially oxidized over an Ag-based
catalyst or metal oxide-based catalyst to selectively produce formalin aqueous solution (FA
concentration 37
�
55
wt
%). In the Ag-based process, MeOH is mixed with an air stream and
fed to a reactor to be converted to FA via partial oxidation and dehydrogenation reactions, as
shown in eqn. 4.4-4.6 and illustrated in Figure 4.11.
CH3OH +1/2O2 → CH2O + H2O (4.4)
CH3OH CO +2H2 (4.5)
CO +1/2O2 CO2 (4.6)
FA(aq.) synthesis takes place at
T>
650
◦C
and near ambient pressure in a kinetically controlled
regime. H
2
O is formed as a by-product. The synthesis was simulated in an adiabatic yield reactor
assuming 98 % conversion of MeOH and a selectivity of 90 % towards FA [108]. The process
concept of this sub-process was presented by Franz et al. [134], which considers the separation
of FA from volatile gases in an absorber column using H
2
O as a washing liquid. Providing
a FA product stream containing about 55
wt
% FA and 45
wt
%H
2
O, this stream should be
concentrated to be further used for the synthesis of longer chain OME. Therefore, this stream is
49
4 OME process description
Figure 4.11:
Simplifed process fow diagram for the production (A) and concentration (B) of
FA(aq.) from MeOH and air based on [4, 109, 134].
fed to a cascade of two evaporators, which split it into two output streams. The target product
is a stream containing about 85
wt
% of FA. This stream is further used for the synthesis of
longer chain OME. A side product stream containing about 10
wt
% of FA is partially used as a
washing liquid for the aforementioned absorber column and partially leaves the sub-process as a
by-product stream.
4.2.3 monomeric FA production
There is no commercial monomeric FA synthesis based on the endothermic dissociation of MeOH
to FA and valuable H
2
(eqn. 4.7), although this route has been investigated since 1960 to identify
selective catalysts [150]. The lack of direct application of the highly reactive monomeric FA
product hinders the market establishment of this production route. In the case of OME synthesis,
this valuable monomeric FA product is important, and thus this sub-process is considered the
“dream reaction“ for the OME value chain. The reaction occurs at high temperatures
>
650
◦C
and requires 85
kJ mol�1
FA. Due to the high reactivity of FA, the retention time is very short
to avoid the formation of the thermodynamically favoured CO, as shown in eqn. 4.8. The main
challenge of this reaction system is to reach high MeOH conversions at a high FA selectivity
without deactivating the catalysts in this strongly reducing H
2
environment. This aspect has
been intensively experimentally investigated in the scientifc community. An overview of various
catalysts is presented by Su et al. [139]. For the implementation in the simulation platform, the
synthesis process described by Sauer et al. [110] combined with the process concept published by
Ouda et al. [128] was adapted. The feedstock MeOH is saturated in a carrier gas and further
dissociated at 900
◦C
to FA and H
2
over an Na-based catalyst, following the complete MeOH
conversion and selectivity experimentally investigated by Sauer et al. [110] of 70 % towards FA
and 30 % towards CO. For the separation of monomeric FA from the reaction products, absorber
columns using mainly MeOH or recycled OME fractions as washing liquids are used. A simplifed
process fow diagram for the production of monomeric FA is illustrated in Figure 4.12.
CH3OH CH2O + H2 (4.7)
CH3OH CO +2H2 (4.8)
50
4 OME process description
Figure 4.12:
Simplifed process fow diagram for the production of monomeric FA from MeOH
based on [4, 110, 128].
4.2.4 OME1 production
Methylal is commercially available, based on MeOH and FA feedstock. The OME
1
sub-process
comprises the conversion of MeOH and FA to OME
1
and is illustrated in Figure 4.13. The
process concept was adapted from Drunsel [129]. The heterogenous catalyzed synthesis takes
place at 60
◦C
and 2
bar
over an acidic catalyst such as A15 in a fxed bed reactor. Besides the
formation of OME
1
, several side reactions take place as shown by eqn. 2.1-2.5. The synthesis
was simulated in an isothermal fxed bed reactor using the kinetic model from Drunsel et al.
[79] as implemented by Bongartz et. al. [108]. The reaction product purifcation takes place
downstream to the reactor in a reactive distillation column. This column is used to overcome
reaction equilibrium restrictions and convert FA almost completely to OME
1
, while separating
H
2
O and MeOH from the azeotropic mixture of OME
1
and MeOH. In a consecutive distillation
column operated at a higher-pressure level, OME
1
is separated from the azeotropic mixture
of OME
1
and MeOH and leaves the distillation column as the bottom product. The distillate
product is recycled to the reactive distillation column.
Figure 4.13:
Simplifed process fow diagram for the production of OME
1
from MeOH and FA(aq.)
based on [4, 129].
4.2.5 Combustion
A combustion sub-process was implemented to use the heating value of the purge streams
to produce process steam, which was utilized in the sub-processes. For the simulation of
the combustion reactions, an adiabatic Gibbs reactor was applied, and excess air was added
accordingly to achieve complete combustion and keep the adiabatic temperature rise below
51
4 OME process description
800
◦C
. The stoichiometric amount of O
2
required for a complete combustion can be estimated
using eqn. 4.9. A simplifed process fow diagram is illustrated in Figure 4.14.
CxHyOz +1/2(2x +1/2y � z)O2 → xCO2 +1/2yH2O (4.9)
Figure 4.14:
Simplifed process fow diagram for the combustion of the purge streams based on
[4].
52
5 Results and Discussion
5.1 Experimental investigation of the OME synthesis
The OME synthesis was investigated for various commercially available catalysts to compare their
activity, selectivity and thermal stability of the synthesis product. Furthermore, the synthesis
was carried out for the feed mixtures OME
1
-TRI as well as MeOH-pFA to consider an anhydrous
and an aqueous reaction system towards OME. In addition, the synthesis products were distilled
in a micro distillation setup to investigate their stability regarding downstream purifcation using
distillation columns. The following sections present the results form [3].
5.1.1 Reaction progress and equilibrium composition
Figure 5.1 and 5.2 illustrate the reaction progress and the equilibrium composition for the OME
synthesis from MeOH-pFA as well as OME
1
-TRI, respectively, at 60
◦C
and 8
bar
over A36.
The analytic results for all investigated catalysts are presented in the appendix section A.1.1.
(a) (b)
Figure 5.1:
OME synthesis from MeOH-pFA over A36 (conditions: pFA/MeOH = 1
.
53
gg
�1
,
A36/(MeOH + pFA) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch) [3]. (a) illustrates the reaction
progress and (b) the equilibrium composition after 24 h.
Besides the progress of the mass fractions of FA, H
2
O, MeOH, OME
1-8
, TRI and MEFO, the
termination time is indicated as vertical line in Figure 5.1a and 5.2a, which is used as an
indicator for the activity of the catalysts and will be discussed in the subsequent section. The
reaction progresses show that the quasi-equilibrium composition is obtained after 3
�
4
h
for
the MeOH-pFA feed mixture and after about 1
h
for the OME
1
-TRI feed mixture. However,
the side product formation shifts the equilibrium composition leading to a slightly diferent
composition after 24
h
. The presence of H
2
O inside the reaction mixture from MeOH-pFA leads
to a reduction of catalyst activity and strongly infuences the selectivity towards OME
3-5
. Due
to the side and intermediate product formation of HF and MG, which are formed in presence
of H
2
O and MeOH as described by eqn. 2.1-2.4, the yield of OME
3-5
reduces from 31
wt
%
54
5 Results and Discussion
(a) (b)
Figure 5.2:
OME synthesis from OME
1
-TRI over A36 (conditions: OME
1
/TRI = 2
.
00
gg
�1
,
A36/(OME
1
+ TRI) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch) [3]. (a) illustrates the reaction
progress and (b) the equilibrium composition after 24 h.
for the OME
1
-TRI feed mixture to 12
wt
% for the MeOH-pFA feed mixture, as illustrated by
the green coloured componentes in Figure 5.1b and 5.2b. Considering the overall composition,
the MeOH-pFA feed mixture leads to a large amount of unreacted feedstock in the equilibrium
composition, while TRI is almost completely converted in the equilibrium composition of the
OME1-TRI feed mixture.
To compare the diferent catalysts, Figure 5.3 illustrates the yield of OME
3-5
over the synthesis
progress for both feed mixtures and all investigated catalysts.
(a) (b)
Figure 5.3:
OME
3-5
yield over the synthesis progress for various catalysts (conditions: pFA/MeOH
=1
.
5
gg
�1
, OME
1
/TRI =2
.
0
gg
�1
, catalyst/reactants = 1
.
0
wt
%, 60
◦C
,8
bar
,
batch) [3]. (a) illustrates the results for the MeOH-pFA feed mixture and (b) illustrates
the results for the OME1-TRI feed mixture.
The yield of OME
3-5
after 24
h
varies between 11
�
14
wt
% for the MeOH-pFA feed mixture and
28
�
34
wt
% for the OME
1
-TRI feed mixture. The progress of the OME
3-5
yield from MeOH-pFA
is much faster for the IER than for the zeolites and Nafon, with Dowex showing a signifcantly
faster reaction than all other IER.
Furthermore, the fnal yield of OME
3-5
using Dowex is higher than for the other catalysts, A15
and H-BEA 25 obtained the lowest yield. These diferences cannot only be explained by the
55
5 Results and Discussion
side product formation, which was especially prominent for the zeolites, however, those show
rather good yields after 24
h
. The results for the MeOH-pFA system are in contrast to those
obtained by Oestreich et al. [24], who reported similar yields between the catalyst systems
and also pronounced side product formations for the zeolites. However, they investigated the
synthesis at higher temperatures of 80 ◦C and ground the zeolites before their application.
Comparing the OME
3-5
yield from MeOH-pFA with OME
1
-TRI, a much faster progress is reached
with the absence of H
2
O, leading to signifcantly higher OME
3-5
yields. For the OME
1
-TRI feed
mixture, the yield also varies between the catalysts and is led by Dowex. For the zeolites, the
OME
3-5
yield decreases after 5
h
due to the strong MEFO formation. This phenomenon will be
discussed in the following section.
Figure 5.4 illustrates the conversion of the reactants MeOH-pFA as well as OME
1
-TRI and the
selectivity towards OME3-5 over various catalysts after 24 h.
(a) (b)
Figure 5.4:
Conversion of the reactants and selectivity towards OME
3-5
for the OME synthesis
from MeOH-pFA and OME
1
-TRI over various catalysts (conditions: pFA/MeOH
=1
.
5
gg
�1
, OME
1
/TRI =2
.
0
gg
�1
, catalyst/reactants = 1
.
0
wt
%, 60
◦C
,8
bar
,
24
h
, batch) [3]. (a) illustrates the results for the MeOH-pFA feed mixture and (b)
the results for the OME1-TRI feed mixture.
The conversion of MeOH after 24
h
in Figure 5.4a shows substantial diferences between the
diferent catalysts, while the conversion of FA is similar. Only for the zeolite H-MFI 90 a higher
conversion of FA was detected, which is mainly attributed to the higher side product formation.
Besides, A15 clearly shows lower and Dowex higher conversions in comparison to the other
catalysts.
Figure 5.4b illustrates the conversions and selectivity towards OME
3-5
for the OME
1
-TRI feed
mixture. Thereby, both reactants led to very similar conversions for all catalysts with a small
diference for the zeolites, which again is a result of their increased side product formation.
The selectivity towards OME
3-5
was found similar between the catalyst systems for both feed
mixtures. However, it was signifcantly lower for the MeOH-pFA mixture because of the presence
of H
2
O and the associated side products. Furthermore, as already indicated by the conversion,
the selectivity of OME
3-5
was slightly lower after 24
h
for the zeolites due to their comparatively
high activity for the MEFO and TRI formation.
56
5 Results and Discussion
5.1.2 Catalyst activity
To determine the activity of the catalysts for the OME syntheses starting from MeOH-pFA as
well as OME
1
-TRI, two indicators were evaluated. The termination time is illustrated in Figure
5.5a and the ratio of the yield of OME
3-5
after 30
min
and 24
hY ∗
is illustrated in Figure
OME3�5
5.5b.
(a) (b)
Figure 5.5:
Termination time (a) and yield
Y ∗
(b) of the OME synthesis from MeOH-
OME3�5
pFA and OME
1
-TRI for various catalysts (conditions: pFA/MeOH = 1
.
5
gg
�1
,
OME1/TRI =2.0 gg
�1, catalyst/reactants = 1.0 wt%, 60 ◦C,8 bar, batch) [3].
For the OME
1
-TRI feed mixture, all catalysts were very active with the shortest termination
time obtained by A15 and A46 followed by H-BEA 25 and Dowex. For the MeOH-pFA feed
mixture, only Dowex was very active followed by the other IER with a clear increase in the
termination time. Oestreich et al. [24] also reported a higher activity for the IER catalysts than
for zeolites for the OME synthesis from MeOH-pFA. No clear tendency was obtained regarding
the acid capacity of the IER as listed in Table 2.2. A36 shows the highest acid capacity but was
found to be less active than A46 and Dowex, with a far lower acid capacity. However, for the
OME
1
-TRI feed mixture, A46 was signifcantly faster, even though their main diference is the
degree of sulfonation which is higher for A36. A36 is sulfonated on the surface and within the
micro pores of the matrix while A46 is only sulfonated on the surface [14]. However, A46 has a
bigger surface than A36. Dowex showed a very high activity for the MeOH-pFA feed mixture
compared to the other IER but a similar activity for the OME
1
-TRI feed mixture. In contrast to
all other catalysts, Dowex showed a higher activity for the MeOH-pFA feed mixture than for the
OME
1
-TRI feed mixture. Therefore, the ring opening of TRI, as described by eqn. 2.9 and the
incorporation into OME as described by eqn. 2.7 are more prominent rate determining steps than
the acetalization reactions from HF to OME as described by eqn. 2.5 and 2.6 and the presence
and formation of the side products HF and MG as described by eqn. 2.1-2.4. Lautenschütz
[25] also reported diferences in the activity of diferent catalysts between an anhydrous and
diferent aqueous reaction systems and explained this with the presence of H
2
O, which leads to
additional side product formations and, therefore, reduces the product selectivity towards OME.
In addition, H
2
O and MeOH inhibit the formation of OME, which was particularly apparent for
H-BEA 25 in comparison to A36.
Regarding the zeolites, H-BEA 25 has a lower Si/Al ratio but is more active than H-MFI 90 for
57
5 Results and Discussion
the OME
1
-TRI feed mixture and less active for the MeOH-pFA feed mixture. Nafon showed
similar activity to the zeolites even though its surface area is signifcantly smaller since it contains
no pores, as presented in Table 2.2. Therefore, smaller Nafon beads could reach the activity
of the IER. The termination time was very similar for A36 and A46 for the MeOH-pFA feed
mixture.
The second indicator for the activity is the yield
Y ∗
which is presented in Figure 5.5b,
OME3�5
it shows higher values for the OME
1
-TRI feed mixture than for the MeOH-pFA feed mixture
for all catalysts except for Dowex. These results match with those from the termination time
of the catalysts. However, the results for
Y ∗
are more precise for evaluating the activ-
OME3�5
ity due to the evaluation of the directly measured composition instead of using linear interpolation.
5.1.3 Side and by-product formation
For the feed mixture OME
1
-TRI, the side products FA, MeOH, H
2
O and MEFO were evaluated,
whereas for the feed mixture MeOH-pFA the side products MEFO and TRI were assessed.
MEFO
Figure 5.6 illustrates the formation of MEFO over the OME synthesis progress for the MeOH-pFA
feed mixture for various catalysts. The dashed lines show the termination times of the respective
catalysts.
Figure 5.6:
MEFO side product formation over the synthesis progress from MeOH-pFA for various
catalysts (conditions: pFA/MeOH = 1
.
5
gg
�1
, catalyst/reactants = 1
.
0
wt
%, 60
◦C
,
8
bar
, batch) [3]. The dashed lines show the termination time of the respective
catalysts.
The highest concentrations of MEFO were obtained with the catalysts H-MFI 90, Dowex and
H-BEA 25. All other catalysts showed very low formations of MEFO for the entire duration,
58
5 Results and Discussion
with concentrations lower than 0
.
1
wt
% MEFO at their respective termination times. Already
before the catalyst was added to the reaction mixture, MEFO concentrations were detected for
all syntheses. After a small increase in the initial phase of about 1
�
3
h
, the concentration of
MEFO stayed approximately constant for the rest of the synthesis without signifcant diferences
between the concentration at the termination time and after 24
h
for all catalysts. This disagrees
with the assumption of the irreversibility of the Tishchenko reaction as described by eqn. 2.12 but
indicates that the reversible esterifcation (eqn. 2.13) is prominent, which requires the presence
of FOAC. Therefore, FOAC was likely to be part of the side products but could not be quantifed
with the applied analysis methods. Following eqn. 2.13, the initial increase of the MEFO curves
was infuenced by the varying MeOH concentration in the reaction mixture, which also stays
approximately constant after the termination time is exceeded. The negligible infuence of the
Tishchenko reaction for all the investigated catalysts is surprising since it leads to signifcant
MEFO concentrations in the OME
1
-TRI feed mixture, as illustrated in Figure 5.7. Furthermore,
Voggenreiter et al. [28] investigated the side product formation for the OME synthesis from
MeOH, FA, OME
1
and H
2
O for the catalyst A46. In contrast to the results obtained with all
catalysts in this work, they reported a steady increase in the MEFO concentration over the
synthesis progress. However, they prepared the feed mixture by dissolving pFA in a solvent using
a base, sodium methoxide or sodium hydroxide to accelerate the process. The MeOH-pFA feed
mixture was prepared at higher temperatures and retention times in this work. This apparently
led to the formation of MEFO even without adding a catalyst. Furthermore, Voggenreiter
et al. [28] investigated the synthesis at lower ratios of the reactants FA and MeOH, with
high concentrations of OME
1
in the reactant mixture and at temperatures between 70
◦C
and
100
◦C
. For their experiment KIN3, almost no OME
1
was present in the reactant mixture and
a small decrease of the MEFO concentration was detected between the initial sample and the
frst reaction sample. However, the intervals between the samples were too big to confrm this
behavior. Comparing the fndings of Voggenreiter et al. [28] with the results from this work,
the Tishchenko reaction seems to be prominent at temperatures exceeding 70
◦C
. However, the
reversible esterifcation was prominent for lower temperatures and reaction mixtures with lower
OME concentrations.
For the OME
1
-TRI feed mixture, the MEFO concentrations were increasing over time for all
catalysts without an indication of reaching an equilibrium composition. Figure 5.7 shows the
synthesis progress until 4
h
to emphasize the initial MEFO formation until the termination
time. For the OME
1
-TRI feed mixture, the irreversible Tishchenko reaction is prominent for
the MEFO formation. The reversible esterifcation is not prominent due to very low MeOH
and H
2
O concentrations in the reaction mixture. Similar to the MeOH-pFA feed mixture, the
zeolites show the strongest MEFO formation. However, in contrast to the high MEFO formation
in the MeOH-pFA feed mixture, MEFO could only be detected for Dowex after exceeding
the termination time. Besides Dowex, the other IER and Nafon also show very low MEFO
concentrations at their respective termination times.
As a conclusion, H-MFI 90 led to very high MEFO concentrations for both feed mixtures already
at the respective termination time. H-BEA 25 showed far lower MEFO concentrations for both
feed mixtures at the respective termination time; however, still exceeding the concentrations
59
5 Results and Discussion
Figure 5.7:
MEFO side product formation over the synthesis progress from OME
1
-TRI for various
catalysts (conditions: OME
1
/TRI =2
.
0
gg
�1
, catalyst/reactants = 1
.
0
wt
%, 60
◦C
,
8
bar
, batch) [3]. The dashed lines show the termination time of the respective
catalysts.
obtained by the other catalysts. Only Dowex showed a higher MEFO concentration for the
MeOH-pFA feed mixture but no MEFO for the OME
1
-TRI feed mixture. The other IER
and Nafon showed similar small MEFO concentrations at their respective termination times.
Unexpectedly, the irreversible Tishchenko reaction was insignifcant for the MeOH-pFA feed
mixture at 60
◦C
. Due to the irreversible Tishchenko reaction, MEFO must be extracted from
the loop inside the OME
3-5
production process to prevent its accumulation. However, due to
a very narrow boiling point curve with OME
1
, the separation from the product mixture can
be expensive [136, 137]. This is an important aspect that should be addressed by extended
experimental investigations for selected catalysts and the infuence of MEFO handling strategies
on the complete process design should be evaluated in further investigations.
TRI
Figure 5.8 illustrates the formation of TRI over the OME synthesis progress for the MeOH-pFA
feed mixture for various catalysts.
The highest concentrations of TRI were obtained with the catalysts A36, A15 and Nafon at their
respective termination times. The catalysts A46 and Dowex showed very low TRI concentrations,
even below the detection limit. The zeolites also led to low TRI concentrations of 0.1 wt% and
below at their respective termination times. In comparison to the MEFO formation, the TRI
curves showed a steady increase. As described by eqn. 2.9-2.11 the formation of TRI from FA,
MG
3
or OME
4
is limited by an equilibrium composition which, however, was not obtained by
any of the investigated catalysts at 60
◦C
until 24
h
. Since TRI also represents a reactant for the
60
5 Results and Discussion
Figure 5.8:
TRI side product formation over the synthesis progress from MeOH-pFA for various
catalysts (conditions: pFA/MeOH = 1
.
5
gg
�1
, catalyst/reactants = 1
.
0
wt
%, 60
◦C
,
8
bar
, batch) [3]. The dashed lines show the termination time of the respective
catalysts.
formation of OME, it does not need to be separated from the loop inside the OME
3-5
production
process. Its presence infuences the reaction kinetics, though, with its concentration limited by
the low equilibrium concentration [27].
Schmitz et al. [27] reported low concentrations of TRI for the OME synthesis from MeOH-pFA
over A46 at diferent temperatures. The concentration of TRI increased with higher concentra-
tions of FA in the feed mixtures and at higher temperatures starting from 70
◦C
. Below 70
◦C
,
no TRI was detected, which agrees with the results in this work for A46. Voggenreiter et al. [28]
also reported low concentrations of TRI in the OME synthesis from MeOH, FA, H
2
O and OME
1
over A46. The amount of TRI increased with rising temperature and FA concentrations in the
feed mixture but mainly did not exceed 1 wt%.
FA, MeOH and H2O
Figure 5.9 illustrates the formation of FA, MeOH and H
2
O over the OME synthesis progress for
the OME1-TRI feed mixture for various catalysts.
The highest concentration of FA and MeOH was obtained with A15, A36 and Dowex. All other
catalysts lead to FA concentrations below 1
wt
% at their respective termination times and even
lower MeOH concentrations.
Considering the curves for the formation of MeOH, all catalysts start at very low concentrations
of less than 0
.
2
wt
%, pass through a maximum and slowly decrease towards a constant concentra-
tion. An exception is Nafon, whose MeOH concentration steadily increases towards a constant
concentration. Considering the reaction network described in a previous section, the formation
of MeOH requires the presence of H
2
O in the reaction mixture. However, Figure 5.9c illustrates
the progress of the H
2
O concentration and except for A36, no H
2
O was detected. Burger et al.
[26] investigated the OME synthesis from OME
1
-TRI over A46 and also did not detect any H
2
O
but low concentrations of MeOH. Lautenschütz [25] analyzed a blank experiment with OME
1
reacting alone in the presence of A36. He reported that the product mixture contained 2
wt
%
61
5 Results and Discussion
(a) (b)
(c)
Figure 5.9:
FA (a), MeOH (b) and H
2
O (c) side product formation over the synthesis progress
from OME
1
-TRI for various catalysts (conditions: OME
1
/TRI = 2
.
0
gg
�1
, cata-
lyst/reactants = 1.0 wt%, 60 ◦C,8 bar, batch) [3].
MeOH and 3
wt
% OME
2
. In a subsequent experiment, he dried OME
1
before adding A36, and
no MeOH or OME
2
could be detected in the product mixture. Therefore, the formation of
MeOH results from traces of H
2
O inside the reaction mixture, which is below the detection limit
of about 0
.
02
wt
%H
2
O. The FA formation is also infuenced by the formation of MeOH from
OME1 and H2O, as described by eqn. 2.5 but additionally by the equilibrium reaction towards
TRI, as described by eqn. 2.9.
Considering the production of OME
3-5
from OME
1
-TRI, the formation of MeOH and, therefore,
the presence of traces of H
2
O inside the feed mixture is challenging for a steady state operation.
Since MeOH and H
2
O would accumulate in the loop inside the OME
3-5
production process, they
need to be separated, which would strongly reduce the beneft of the anhydrous OME reaction
system compared to the aqueous OME reaction system. Alternatively, the reactants need to be
intensively dried before application. This would refect on the production costs but entail the
advantage of a signifcantly simplifed OME3-5 product separation and purifcation.
62
5 Results and Discussion
5.1.4 Thermal stability of the synthesis products
Using IER to synthesize OME can lead to leaching of the active acid groups of the catalyst
into the reaction product mixture [32]. This was also emphasized by Baranowski et al. [105],
who concluded the main drawbacks of IER to be the low thermal stability and the leaching
of active species into the synthesis product when using polar solvents. Acid IER are mainly
synthesized by copolymerization of styrene and divinylbenzene with a macroreticular matrix and
functionalized with sulphuric acid [151]. Therefore, the functional (-SO
3
H) groups can leach into
the reaction mixture. Fink [152] investigated the stability of A15 and A36 for the OME synthesis
from MeOH-pFA at 65
�
70
◦C
and reported that 0
.
4
mol
% and 0
.
7
mol
% of the sulphur content
from the sulphonic acid group was dissolved out the catalyst after 3
�
5
h
. As a result, she
concluded that IER are generally not suitable for the OME
3-5
production. Furthermore, before
the thermal separation of the OME synthesis product mixture using IER, Lautenschütz [25]
neutralized the mixture with IER III but did not mention its necessity.
Active species inside the OME synthesis product enable the reactions towards and between
diferent OME as described by eqn. 2.5-2.8, but also side product formations as described by eqn.
2.9-2.16, outside the reactor unit. Due to the comparatively slow kinetics of these reactions, traces
of active species will not show a signifcant infuence on the product composition at moderate
temperatures. However, considering a thermal separation of the OME synthesis product to purify
the desired OME
3-5
fraction, high temperatures of about 200
◦C
[1] are required, which strongly
accelerate the reactions. Furthermore, due to the separation of the more volatile components
from the OME synthesis product the reaction equilibrium of eqn. 2.5-2.8 is disturbed and the
direction of the reaction will reverse. This results in the formation of MeOH and HF as well
as shorter chain OME and FA following the reactions describes by eqn. 2.5-2.7. Furthermore,
the transacetalization reactions as described by eqn. 2.8 will form shorter chain OME and even
longer chain OME. Depending on the reaction kinetics, temperature level and duration, the
composition of the bottom product of the distillation column varies and the amount of bottom
product will reduce, if active species are present. This reduces the originally produced amount of
the fnal OME
3-5
from the synthesis and needs to be prevented. A neutralization step contacting
the free active acid groups with alkaline groups (OH-) after the reactor can neutralize active
species and enable a stable thermal separation of the OME synthesis product.
To investigate the necessity of a neutralization step, the OME synthesis products were distilled
and the composition after the distillation was compared to the composition before the distillation.
Figure 5.10 illustrates the results of the distillation with the flled bars showing the composition
before the distillation, and the striped bars showing the composition after the distillation, which
is the sum of the distillate and the bottom product.
The synthesis product from MeOH-pFA illustrated in Figure 5.10a shows similar compositions to
the product after the distillation for all catalysts. Only the distillation product of the zeolites
shows more signifcant diferences for FA, MeOH and OME
1
. Unexpectedly, the concentration of
OME
1
reduced during the distillation of the OME synthesis product from H-MFI 90 and led
to higher concentrations of MeOH and FA. Due to the lower volatility of OME
1
, active species
inside the synthesis product should increase the total amount of OME
1
, which is used inside the
63
5 Results and Discussion
(a)
(b)
Figure 5.10:
Synthesis product composition and added up distillation product composition of
the OME synthesis from MeOH-pFA (a) and OME
1
-TRI (b) for various catalysts
(conditions synthesis: pFA/MeOH = 1
.
5
gg
�1
, OME
1
/TRI = 2
.
0
gg
�1
, cata-
lyst/reactants = 1
.
0
wt
%, 60
◦C
,8
bar
, batch; conditions distillation: 30
�
50
g
synthesis product, TOil = 60 � 100 ◦C stepwise, 5 h, batch) [3].
reactive distillation of the OME
1
production process [129]. However, the reverse acetalization
reactions described by eqn. 2.5 and 2.6 might be accelerated faster than the reverse chain
propagation reaction described by eqn. 2.7. Therefore, due to kinetic limitations, this can lead
to higher concentrations of MeOH and FA. H-BEA 25, on the other hand, leads to a reduction of
FA, an increase of the MeOH concentration and a slight increase of the longer chain OME and
by-products TRI and MEFO. A36 and A46 also obtained minor diferences for FA, MeOH, H
2
O
and OME
1
. In contrast to the other catalysts, the bottom product of A46 solidifed at room
temperature and was dissolved in MeOH for analysis. Considering the bottom composition, the
FA, H
2
O and OME
≥4
concentration is slightly higher for A46 compared to the other catalyst
products, and the MeOH concentration is somewhat lower, which might exceed the solubility
limits of FA and longer chain OME. The bottom product compositions after the distillation are
presented in the Appendix in section A.1.2.
Considering the result of the distillations of the synthesis product from OME
1
-TRI as illustrated
in Figure 5.10b, more signifcant diferences were obtained. Only A36, H-BEA 25 and Nafon show
64
5 Results and Discussion
very similar results after the distillation. The distillation of the OME synthesis product of A46
and H-MFI 90 led to a reduction of the short chain OME
1-2
and an increase of the longer chain
OME
≥4
, which indicates the chain propagation coupled with the transacetalization reactions,
as described by eqn. 2.7 and 2.8. However, the bottom product of A46 also solidifed at room
temperature for the OME
1
-TRI based synthesis product and was dissolved in MeOH for analysis.
Similar to the bottom product of the MeOH-pFA synthesis product, the concentration of longer
chain OME
≥4
increased for A46, which is illustrated in the Appendix in Figure A.4. Dowex,
on the other hand, leads to a substantial increase of the OME
1
concentration and decreased
concentrations of OME
≥2
due to the reverse chain propagation described by eqn. 2.7. The FA
concentration did not change as expected according to eqn. 2.7, however, a complete condensation
of the gaseous FA without solidifcation is challenging, especially without the presence of H
2
O
and MeOH. Therefore, the measured concentration of FA inside the bottom and distillate product
can difer signifcantly from the actual amount of FA inside the setup. The bottom product of
the distillation of the OME synthesis product over Dowex also solidifed and was dissolved in
MeOH for the analysis.
A neutralization of the synthesis products using IER III led to the expected composition after
the distillation and prevented the bottom product solidifcation.
Overall, the results of this section indicate that some catalysts lead to OME synthesis products
that are thermally unstable and consequently further react in thermal separation process steps.
Other catalysts show stable product behavior, such as Dowex and Nafon for the MeOH-pFA feed
mixture and A36, H-BEA 25 and Nafon for the OME
1
-TRI feed mixture. However, the tests
were only conducted at an oil bath temperature of up to 100
◦C
. For the distillation separation
of the OME synthesis product for the purifcation of OME
3-5
, temperatures of about 200
◦C
are
required [1]. This temperature increase coupled with the residence time inside the distillation
column would lead to a substantial acceleration of the reaction kinetics and consequently reduce
the OME
3-5
product amount. Therefore, the thermal stability of the OME synthesis product
should be tested at conditions close to the operational conditions inside the distillation column to
decide the consideration of a neutralization step after the OME synthesis. Furthermore, the OME
synthesis products were prepared with fresh catalysts, the thermal instability of the synthesis
product might also be an initial phenomenon that might reduce with increasing time on stream
after washing out instable acid groups.
Besides IER III other heterogeneous and homogeneous alkaline beds and solutions might be
feasible for the neutralization of the OME synthesis product, such as alkaline IER, CaO, MgO
or alkaline loaded active carbon. For further investigations not only process unit specifc
performance indicators such as activity, side product formation, long-term stability and regen-
eration should be considered, but also the infuence on the overall process design and performance.
65
5 Results and Discussion
5.2 Experimental demonstration of the main COMET process units
In the following sections the experimental results of the demonstration of the main COMET
process units are presented and discussed. A simplifed process fow diagram is illustrated in
Figure 4.7. All investigations were carried out using state-of-the-art experimental setups. The
FA concentration units and the OME synthesis unit were interconnected. The products were
collected and further processed in the distillation units consequently. The analytic results are
presented in the appendix in section A.2. The following sections present the results from the
submitted publication, see section 1.2.
5.2.1 OME synthesis
The continuous OME synthesis was investigated experimentally by feeding pure OME
1
and a
concentrated FA(aq.) solution in a fxed bed reactor with 2
.
7
�
3
.
5
Lh
�1
at about 90
◦C
and
10
bar
over the catalyst A46. Figure 5.11 illustrates the composition of the feed mixture (F),
the simulated equilibrium product composition (P-Sim), two preliminary experimental products
(P1-2-Exp) from the starting phase and the three product barrels (P3-5-Exp) which contained
about 250
kg
of the OME synthesis product. The analytic results are presented in the appendix
in section A.2.1.
Figure 5.11:
OME synthesis from OME
1
and concentrated FA(aq.) over A46 (conditions: concen-
trated FA(aq.) with 85
�
89
wt
% FA, (concentrated FA(aq.))/OME
1
=0
.
6
gg
�1
,
A46/(OME
1
+concentrated FA(aq.)) = 0
.
34
gh g�1
, approx. 3
Lh
�1
, 90
◦C
,
10
bar
, fxed bed reactor) [4]. F represents the feed composition. P-Sim, P1-5-Exp
represent the product composition of the simulated equilibrium, two experimental
preliminary products P1-2 from the starting phase and the three product barrels
P3-5, respectively.
The compositions of the three product barrels (P3-5-Exp) show a good agreement with the
simulated equilibrium composition (P-Sim). Comparing the three product compositions among
each other, a small shift towards longer chain OME
≥ 3
and FA with increasing time on stream was
observed. This is mainly a result of the slightly fuctuating FA concentration in the concentrated
FA(aq.) feed stream (85
�
88
wt
% FA) and the feed stream fowrates. The OME
1
fowrate was
66
5 Results and Discussion
regulated to meet a constant ratio between OME
1
and FA, while the FA fowrate was regulated
to stabilize the level of the small storage between the second thin flm evaporator and the OME
synthesis sub-process. Besides a good agreement of the chemical equilibrium between simulation
and experiment, the reaction kinetics of the OME synthesis were predicted much faster than
experimentally investigated. The simulation predicted the chemical equilibrium at a weight
hourly space velocity (WHSV, feed mass fow rate in relation to the amount of catalyst) of about
70
h�1
. However, the experiments were carried out at a WHSV of approximately 16
h�1
and
3
h�1
, whereby only the lower WHSV was sufcient to obtain chemical equilibrium, as presented
by P3-5-Exp in Figure 5.11. The WHSV of 16 h�1 led to higher amounts of unreacted FA, low
concentrations of OME
≥ 3
and, therefore, solidifcation of the synthesis product after cooling and
without adding MeOH for stabilization. The results of the synthesis product composition at the
WHSV of 16
h�1
are presented by P1-2 Exp in Figure 5.11. The kinetic model from Schmitz et
al. [80] was used for the simulation, which was initially regressed on experimental results of the
OME synthesis from MeOH and FA with partly higher concentrations of H
2
O or OME inside.
Therefore, the feed mixture already contained high concentrations of HF which can directly react
to OME, as described by eqn. 2.5 and 2.6. Furthermore, the model was based on the assumption
that the reactions towards HF and MG, as described by eqn. 2.1-2.4 are in equilibrium at all
retention times since their kinetics are much faster than the kinetics of the formation of OME. In
the COMET process this assumption is not met. The concentration of MeOH in the feed is very
low because OME
1
was used as methyl group supplier instead. Therefore, FA is bound mainly in
MG which need to depolymerize to be converted to OME, as described in section 4.1.7. This
is the limiting step for the reaction kinetics of the COMET process but not signifcant for the
OME synthesis based on MeOH and FA. Therefore, the kinetic model [80] is a suitable basis but
needs to be further extended to realistically describe the reaction progress of other feed mixtures,
which is required to correctly dimension the reactor unit.
Besides the main components, small fractions of the side products MEFO, TRI and tetroxane
were detected in the product barrels P3-5 of about 0
.
1
wt
%, 0
.
6
wt
% and 0
.
1
wt
%, respectively.
For P1-2 concentrations of about 0
.
1
wt
%, 0
.
1
wt
% and 0
.
03
wt
% were obtained. Therefore,
these concentrations strongly depend on the WHSV, which is a matter of investigation for high
yields of OME3-5 at low concentrations of the side product.
5.2.2 Synthesis product neutralization
Before the separation of the OME synthesis products P3-5 in the continuous distillation setup,
the thermal stability was tested with a similar procedure to the investigations discussed in section
3.1.7. The pre-tests were conducted in a micro distillation setup using about 50
mL
of P3-5
with a stepwise increase of the reboiler temperature. The distillation led to a solidifcation of
the bottom products for all three samples. However, while P3 solidifed at about 130
◦C
, P4
only solidifed at about 170
◦C
and P5 solidifed only after heating up to 170
◦C
and cooling
down the bottom product. Additionally, the electric conductivity was measured for the three
samples with a reduction from 6
.
3
µS cm�1
for P3, 3
.
2
µS cm�1
for P4 and 2
.
1
µS cm�1
for
67
5 Results and Discussion
P5. This indicates that the reduced thermal stability of the OME synthesis product is an initial
phenomenon. Nevertheless, for the thermal separation of P3-5 a neutralization was necessary.
Further pre-tests were conducted with diferent retention times of the IER III in the OME product
mixtures for neutralisation. The results indicated that a WHSV of about 12
h�1
was sufcient
to neutralize the OME synthesis product at ambient temperature. This WHSV includes a high
safety margin and can probably be increased, especially for the OME synthesis product after
higher times on stream of the catalyst. Furthermore, the results showed a connection between
the thermal stability of the synthesis product as verifed by the micro distillation and the electric
conductivity. Below an electric conductivity of 0
.
2
µS cm�1
the synthesis product was thermally
stable, and the distillation did not change the product composition. Above 1
.
0
µS cm�1
changes
in the composition were detected. Finally, the OME synthesis product of the three product
barrels P3-5 was neutralized at ambient temperature in a fxed bed of IER III at a WHSV of
about 12
h�1
and about 500
g
IER III. A deactivation of the IER III with increasing time on
stream was not observed for the OME synthesis product, of about 200
kg
which was continuously
neutralized in the fxed bed.
During the investigation of the OME synthesis for about 80
h
on stream, a stable catalytic
activity was noticed. Also under reactive distillation conditions, the catalyst performance did
not show an obvious deactivation for about 600
h
on stream. However, further investigations are
required to verify if the changing thermal stability is an initial phenomenon of the catalysts time
on stream. In addition to the impact on the process design, the cause of this behavior should
be investigated. It might only be the leaching of the catalyst as emphasized by Fink et al. [32,
152] and Baranowski et al. [105], but it could also be infuenced by the side product formation,
especially FOAC, which was not analyzed in this work but reported in the literature [28].
5.2.3 Synthesis product separation in CO-1
After the neutralization of the OME synthesis products P3-5 the distillation of the volatile
components FA, H
2
O, MeOH and OME
1-2
from OME
≥ 3
was investigated in the distillation
setup. The results illustrated in Figure 5.12 are an exemplary result of the continuous distillation
experiment and show that the separation between OME
2
and OME
3
was successfully realized
and that FA, H
2
O and MeOH can be separated from OME
≥ 3
. The side products MEFO and
TRI are also separated from the bottom product, but the side product tetroxan has a higher
boiling point than OME
3
and stays in the bottom product. The analytic results are presented in
the appendix in section A.2.2.
The distillation setup was operated at a feed temperature of 87
◦C
, a condensation temperature
of 85
◦C
and a reboiler temperature of 175
◦C
. The distillate to feed ratio was about 0
.
81 and the
time-based refux ratio was varied as a controlled variable between 0
.
5
�
2
s s�1
(time controlled)
to achieve a constant condensation temperature. Figure 5.12 shows that OME
2
was completely
separated to the distillate product. However, also a small fraction of OME
3
went to the distillate
product, which was about 14 % of the feed amount of OME
3
. Besides OME
≥ 3
, traces of FA,
68
5 Results and Discussion
Figure 5.12:
CO-1, OME synthesis product separation (conditions: 2
Lh
�1
, refux/distillate
=0
.
5
�
2
ss
�1
, distillate/feed = 81
wt
%, Montz 750 structured packing, 85
�
175
◦C
,
ambient pressure) [4]. The values describe the mass fractions of the feed mixture,
here P5-Exp as presented in Figure 5.11, the distillate product and bottom product.
H
2
O and MeOH were detected in the bottom product which were mainly below 0
.
6
wt
%. MEFO
was not detected inside the bottom product.
Regarding the continuous operation of the distillation setup, an increasing precipitation of FA
inside the condenser was challenging in the initial phase but could be prevented by increasing the
temperature of the cooling fuid to above 25
◦C
. However, as a result the temperature diference
decreased between the cooling fuid and the boiling points of the most volatile components
MEFO, the azeotropic mixture of OME
1
and MeOH as well as OME
1
. Thus, the surface area of
the condenser was relatively small to obtain a complete condensation and small fraction of the
most volatile components accumulated in a cool trap. As a result, the ratio of OME
1
to OME
2
in the feed mixture P5-Exp difers from the ratio of OME1 to OME2 in the distillate product.
5.2.4 Reactive distillation in CO-2
After the separation of the OME synthesis product in CO-1, the distillate product of CO-1 was
separated and converted in a reactive distillation column.
An exemplary result of the continuous reactive distillation experiment is illustrated in Figure
5.13. The distillate and bottom product compositions show that the targets of the reactive
distillation column were obtained. OME
≥ 2
were converted to OME
1
and FA, the composition of
the distillate product is the azeotropic mixture of OME
1
and MeOH and the bottom product
contains mainly FA and H
2
O. Regarding the bottom product composition, besides the desired
range of FA and H
2
O, only small concentrations of MeOH of about 0
.
3
wt
% were detected.
Furthermore, traces of OME
1-6
were detected with concentrations far below 0
.
1
wt
%. However,
due to the high H
2
O and FA content in the bottom product, the quantifcation of traces is
complex and further complicated because the bottom product solidifes fast if not heated or
diluted. The analytic results are presented in the appendix in section A.2.3.
69
5 Results and Discussion
Figure 5.13:
CO-2, reactive distillation of the distillate product of CO-1 over A46 (conditions:
A46/(feed stream) = 0
.
35
gh g�1
,1
L h�1
, distillate/feed = 63
wt
%, Montz 750
structured packing, 45
�
104
◦C
, ambient pressure) [4]. The values describe the
mass fractions of the feed mixture, the distillate product and bottom product.
The distillation setup was operated at ambient pressure, a condensation temperature of 45
◦C
and reboiler temperature of 104
◦C
with a distillate to feed ratio of about 0
.
63. The technical
feasibility of the reactive distillation column was demonstrated for a long duration of around
600 h on stream.
The results confrm that the reactive distillation column is a feasible instrument for the separation
of H
2
O from the loop of the OME
3-5
production and that an almost complete conversion of
MeOH can be achieved. Furthermore, the results indicate, that the variation of the amount of
MeOH in the feed mixture to the reactive distillation column can be used to set the amount
of OME
1
produced as the distillate product. This defnes if OME
1
can be extracted from the
process as an additional side product.
5.2.5 Product separation in CO-3
The results of the separation of the fnal product mixture OME
3-5
from the bottom product
of the distillation column CO-1 are illustrated in Figure 5.14 as an exemplary result of the
continuous distillation experiment. The target was a cut between OME
5
and OME
6
. Thereby, a
signifcant amount of OME
5
stayed in the bottom product, which, however, can be separated to
the distillate product by increasing the reboiler temperature or reducing the operational pressure.
Regarding the distillate product, besides OME
3-5
small fractions of OME
6
of 0
.
2
wt
%, tetroxan
of 0
.
7
wt
% and traces of FA and H
2
O were detected. The concentration of tetroxan is mainly
a result of the retention time, temperature, and selection of catalyst in the reactor, which can
be improved to reduce the side product formation. However, the pre-standard DIN/TS 51699
does not limit the concentration of tetroxan. The concentration of TRI is limited to 0
.
1
wt
%
and was detected smaller than 0
.
01
wt
%. The fnal OME
3-5
product from the COMET process
was compatible with the pre-standard DIN/TS 51699. The analytic results are presented in the
appendix in section A.2.4.
70
5 Results and Discussion
Figure 5.14:
CO-3, product separation (conditions: 5
.
5
Lh
�1
, distillate/feed = 82
wt
%, Montz
750 structured packing, 100
�
210
◦C
, 200
mbar
) [4]. The values describe the mass
fractions of the feed mixture, here the CO-1 bottom product, the distillate product
and bottom product.
Due to the solidifcation of the bottom product at ambient temperature, it was diluted in THF
with a ratio of 1 : 10
gg
�1
to enable the GC analysis. However, this also increases the detection
limits and accuracy of the analysis. To liquify the bottom product it can also be heated up.
At 80
◦C
, the bottom product is already completely liquid, which enables its recycling to the
reactor, as illustrated in Figure 4.7.
The distillation setup was operated at a condensation temperature of 140
◦C
and a reboiler
temperature of 210
◦C
. The high distillate temperature was a result of the high feed fow rate
and the limited area for condensation. As a result, a complete condensation was not obtained
and a small fraction of OME
3
accumulated in a cool trap. In contrast to the other distillation
experiments the operation pressure was reduced to 200
mbar
to reduce the reboiler temperature
for the separation between OME5 and OME6. The distillate to feed ratio was about 0.82.
71
5 Results and Discussion
5.3 Process simulation
The following sections present the results from the submitted publication, see section 1.2, and [1].
5.3.1 Process description
The process chain for the production of OME
3-5
starts from the production of MeOH, which is
synthesized from H
2
and CO
2
at 250
◦C
and 70
bar
in the gas phase [149]. For the purifcation
the reactor product enters a frst fash unit at 65
◦C
and 65
bar
, a second fash unit at 66
◦C
and 1
bar
and a distillation column operated at 1
bar
, to separate MeOH and H
2
O, see Figure
5.15 [107].
5.3.1.1 Process route P1 - MeOH and FA(aq.) feedstock
After the synthesis and purifcation of MeOH, FA(aq.) is produced and used downstream for the
synthesis of OME3-5, as illustreated in Figure 5.15.
Figure 5.15:
Simplifed process fow diagram of P1 for the production of OME
3-5
from H
2
and
CO2 via the intermediate production of MeOH and FA(aq.) [1].
One part of the MeOH intermediate product is mixed with air and H
2
O and sent to the gas
phase synthesis of FA via partial oxidation and dehydrogenation over a silver catalyst at 650
◦C
and 1
bar
, following eqn. 4.4. Subsequently, FA is separated from the gas stream in an absorber
column using H
2
O as a washing solution [108]. Due to the high H
2
O concentration of about
45
wt
%, the product stream is concentrated in a cascade of two evaporators to a FA rich stream
containing about 86
wt
% FA and a H
2
O rich stream containing about 10
wt
% FA, which is
partly used as washing solution in the FA absorber column [13]. The FA rich stream is mixed
with the second part of the MeOH product and used as a feed stream for the production of
OME
3-5
. The synthesis of longer chain OME in the liquid phase at 80
◦C
and 2
bar
in presence
of an acidic heterogeneous catalyst, such as A46, leads to a variety of side products which are
72
5 Results and Discussion
separated in two distillation columns and a membrane unit, which separates the side product
H
2
O [18]. The frst distillation column operates at about 1
bar
and provides a bottom stream at
186
◦C
comprising mainly OME
3-10
and a distillate stream at 64
◦C
, which contains the rest of
the components and a slip of OME
3
, which amounts for about 42 % of the OME
3
in the feed
stream. The distillate stream is send to the membrane unit to separate H
2
O from the loop, and
recycled to the synthesis of OME [92]. The bottom stream is send to the second distillation
column operating at 0
.
078
bar
which provides the main product stream containing OME
3-5
in
the distillate at 80
◦C
and a bottom stream of OME
6-10
at 196
◦C
, which is recycled back to the
OME synthesis. In comparison to the feed streams comprising MeOH and concentrated FA(aq.),
the mass fow of the recycled streams containing OME
1-3,6-10
, MeOH, FA and a rest of H
2
O is
about 5.4 times larger.
5.3.1.2 Process route P2 - MeOH and monomeric FA feedstock
After the synthesis and purifcation of MeOH, monomeric FA is produced and used downstream
for the synthesis of OME3-5, as illustrated in Figure 5.16.
Figure 5.16:
Simplifed process fow diagram of P2 for the production of OME
3-5
from H
2
and
CO2 via the intermediate production of MeOH and monomeric FA [1].
One part of the MeOH intermediate product is saturated in N
2
to a concentration of 15
vol
% to
synthesize FA via dehydrogenation over Na
2
CO
3
or NaAlO
3
at 900
◦C
and 2
bar
. FA is separated
from the gas stream in an absorber column using the second part of the MeOH intermediate
product as washing solution [110, 139]. This separation was adopted and adjusted from the
FA(aq.) sub-process, which uses H
2
O as a washing solution instead. H
2
O is not a suitable
washing solution for the absorbtion of monomeric FA for the production of OME, since it needs
to be separated from the loop of the OME sub-process and reduces the selectivity towards
OME
3-5
. Using MeOH instead of H
2
O as a washing solution is a new concept and was not
experimentally validated before. However, due to the similar reaction system between MeOH and
FA in comparison to FA and H
2
O (see eqn. 2.1-2.4), it is expected to yield a satisfying separation
73
5 Results and Discussion
at adjusted operation parameters. The product mixture of the absorber column contains about
63
wt
% FA and is used as a feed stream for the production of OME
3-5
. The conditions of the
process units inside the OME sub-process are similar to the P1 process. In comparison to P1,
the amount of OME
3
leaving the frst distillation column to the distillate product is reduced to
about 27 % of the OME
3
content of the feed mixture. Moreover, the ratio of the mass fows of
the recycled streams to the feed stream reduces to 4.4.
5.3.1.3 Process route P3 - OME1 and FA(aq.) feedstock
After the synthesis and purifcation of MeOH, FA(aq.) and OME
1
are produced and used
downstream for the synthesis of OME3-5, as illustreated in Figure 5.17.
Figure 5.17:
Simplifed process fow diagram of P3 for the production of OME
3-5
from H
2
and
CO2 via the intermediate production of MeOH, FA(aq.) and OME1 [1].
One part of the MeOH intermediate product is used to produce FA following the FA(aq.) sub-
process, as described for P1 in section 5.3.1.1. Parts of the bottom product stream of the FA
absorber column is mixed with the second part of the MeOH product and used to produce OME
1
.
OME
1
is synthesized over an acidic heterogeneous catalyst, such as A15 at 60
◦C
and 2
bar
and purifed using a series of a reactive distillation column with catalytic zones and a second
distillation column [108]. The reactive distillation column is operated at 1
bar
and produces a
distillate product stream containing the azeotropic mixture of OME
1
and MeOH with about
94
wt
% OME
1
at 40
◦C
, a bottom product stream containing mainly H
2
O at 98
◦C
and a
gaseous side product stream below the catalytic zone containing 84
wt
% MeOH at 67
◦C
, which is
recycled back to the OME
1
reactor. The second distillation column splits the azeotropic mixture
of OME
1
and MeOH at a higher pressure of 4
bar
into a distillate product stream containing the
new azeotropic composition of about 91
wt
% OME
1
and MeOH at 85
◦C
and an almost pure
OME
1
bottom product stream at 88
◦C
. The OME
1
product stream is mixed with the FA rich
stream and used as a feed stream for the production of OME
3-5
. The synthesis and purifcation of
74
5 Results and Discussion
OME
3-5
is similar to P1. In comparison to P1, the amount of OME
3
leaving the frst distillation
column in the distillate product stream is reduced to 37 %. Moreover, in comparison to the feed
streams the mass fow of the recycled streams is 4.8 times larger.
5.3.1.4 Process route P4 - OME1 and monomeric FA feedstock
After the synthesis and purifcation of MeOH, monomeric FA and OME
1
are produced and used
downstream for the synthesis of OME3-5, as illustreated in Figure 5.18.
Figure 5.18:
Simplifed process fow diagram of P4 for the production of OME
3-5
from H
2
and
CO2 via the intermediate production of MeOH, monomeric FA and OME1 [1].
P4 is a combination of P2 and P3 and uses one part of the MeOH intermediate product to
produce FA, following the monomeric FA sub-process, as described for P2 in section 5.3.1.2.
One part of the product stream of the FA reactor is sent to an absorber column in which the
second part of the MeOH product is used as a washing solution. The bottom product stream
containing about 68
wt
% MeOH is used for the production of OME
1
as described for P3 in
section 5.3.1.3. The absorption of the second part of the product stream of the FA reactor uses a
recycled stream of the OME
3-5
sub-process, which mainly contains OME
1-3
. This absorption
concept should be experimentally investigated, since the solubility of monomeric FA in OME
mixtures not containing MeOH or H
2
O is still unknown. However, substituting the OME mixture
with MeOH or H
2
O is not a suitable option, since their presence in the synthesis of longer chain
OME would increase the side product formation and, therefore, reduce the selectivity towards
OME
3-5
. The product stream containing FA and a mixture of OME
1-3
is mixed with the OME
1
product stream and used as a feed stream for the production of OME
3-5
. The sub-processes
of OME
1
and OME
3-5
are similar to P3. The main diference is the amount of H
2
O separated
in the sub-process of the OME
3-5
production. Starting from OME
1
and monomeric FA, only
very small amounts of H
2
O and MeOH enter this sub-process in form of impurities. However, to
prevent its accumulation, H
2
O needs to be separated form the loop. In comparison to P1 only
75
5 Results and Discussion
31 % of OME
3
leaves the frst distillation column in the distillate product stream. The ratio
between the mass fow of the feed stream and the recycled streams is reduced to 3.1.
5.3.1.5 COMET process
Stream compositions and conditions of the COMET process simulation in Aspen Plus
®
are listed
in Table 5.1, following the stream numbering of Figure 4.7. Stream compositions and conditions
for all sub-processes are presented in the appendix in section A.3.4.
Table 5.1:
Stream compositions and conditions of the COMET process presented in Figure 4.7
[4].
Overall mass fractions
Stream 1 2 3 4 5 6 7 8 9 10 11 12 13 14
T in ◦C 64.9 30 90 30 90.4 90 90 81.5 81 41.5 117.4 200.5 86.6 194.9
p in bar 1 1 0.3 1 10.3 10 10.1 1.8 1.8 1 1 1.8 0.07 0.07
m in kg h�1 18509 2203 14753 7870 22957 66666 66666 51796 5288 41330 15753 14871 12490 2380
FA 0.502 0.142 0.184 0 0.88 0.303 0.186 0.239 0 0 0.727 0 0 0
H2O 0.491 0.778 0.796 1 0.12 0.042 0.022 0.028 0 0.002 0.268 0 0 0
MeOH 0.007 0.08 0.02 0 0 0.028 0.1 0.129 1 0.045 0.005 0 0 0
OME1 0 0 0 0 0 0.591 0.276 0.356 0 0.953 0 0 0 0
OME2 0 0 0 0 0 0 0.179 0.23 0 0 0 0 0 0
OME3 0 0 0 0 0 0 0.107 0.017 0 0 0 0.419 0.499 0
OME4 0 0 0 0 0 0 0.061 0 0 0 0 0.271 0.323 0
OME5 0 0 0 0 0 0 0.033 0 0 0 0 0.149 0.177 0
OME6 0 0 0 0 0 0.018 0.018 0 0 0 0 0.08 0 0.496
OME7 0 0 0 0 0 0.009 0.009 0 0 0 0 0.042 0 0.262
OME8 0 0 0 0 0 0.005 0.005 0 0 0 0 0.022 0 0.136
OME9 0 0 0 0 0 0.002 0.002 0 0 0 0 0.011 0 0.07
OME10 0 0 0 0 0 0.001 0.001 0 0 0 0 0.006 0 0.036
The feedstock (stream 1) containing about 50
wt
% FA and 49
wt
%H
2
O is mixed with the
distillate of the second evaporator E-2 and the bottom of the third evaporator E-3. The
mixture is concentrated in a cascade of two evaporators E-1 and E-2 operated at 400 and
500
mbar
respectively and low retention times. The pressure levels were selected to obtain similar
evaporation and condensation temperatures as experimentally verifed. However, in practice
the pressure level might be lower to achieve the desired concentrations. This is a result of the
simplifed modelling of the evaporators which require more detailed considerations of the reaction
kinetics of eqn. 2.1-2.4 as recently introduced by Tönges and Burger [153]. The FA concentration
is similar to the production of pFA and generates a concentrated FA solution containing about
88
wt
% FA (stream 5) and a solution containing about 18
wt
% FA (stream 3). Stream 3 is split
to be used as a washing liquid for the FA absorber column and to be purifed in the distillation
column CO-4 operated at 5
.
5
bar
to pure H
2
O (stream 4) (
<
200
ppm
FA) and a concentrated
FA solution with 44
wt
% FA. To prevent the accumulation of MeOH and other impurities in
the loop, the concentrated FA solution is sent to another evaporator E-3 operated at ambient
pressure. This prepares a by-product of the COMET process with a higher MeOH concentration
(stream 2) and a FA solution with a similar composition to the FA feedstock, which is recycled to
the evaporator cascade. The by-product (stream 2) has a low FA concentration of about 14
wt
%.
Furthermore, its mass fow is about 17
.
6 % of the mass fow of the target OME
3-5
product. This
76
5 Results and Discussion
is similar to alternative OME
3-5
production processes using FA(aq.) solution as an intermediate
product [1].
The concentrated FA product (stream 5) is pressurized to about 10
bar
, then mixed with the
recycle streams (stream 10 and stream 14) and converted to OME in a fxed bed reactor at
about 10
bar
and 90
◦C
, over A46 catalyst, as used for the experimental demonstration. The
reactor product contains about 20
wt
% OME
3-5
, which is relatively high in comparison to the
process based on MeOH and FA(aq.) with 0 to 15
wt
% OME
3-5
in the reactor product [109],
as presented in Table 5.5. The reactor product is purifed in a frst distillation column CO-1
operated at a slight overpressure of 1
.
8
bar
, where OME
≥ 3
are separated from FA, MeOH,
H
2
O, OME
1-2
and a small fraction of OME
3
. The slight overpressure improves the separation
efciency and reduces the losses of OME
3
to the distillate product (stream 8) to about 12 %.
The FA concentration of the bottom product (stream 12) is reduced to about 100
ppm
. In the
third distillation column CO-3 operated at 70
mbar
, the main product OME
3-5
(stream 12) is
extracted from the process with about 50
wt
% OME
3
, 32
wt
% OME
4
, 18
wt
% OME
5
and traces
of FA and H
2
O in compliance with the pre-standard DIN/TS 51699 specifcations. The distillate
product (stream 8) of the frst distillation column CO-1 is mixed with MeOH (stream 9) from
the MeOH sub-process and introduced to the reactive distillation column CO-2. This column is
operated at ambient pressure. The selection of the pressure level is a compromise between the
condensation temperature of the distillate, the reaction kinetics on the catalytic trays and the
composition of the azeotropic mixture of OME
1
and MeOH in the distillate. A pressure reduction
would favorably improve the azeotropic composition to higher OME
1
concentrations. However, it
would also lead to a reduction of the condenser temperature below 41
◦C
which can lead to more
expensive cooling utilities and decelerate the reaction kinetics on the catalytic trays. Increased
pressure levels would beneft from higher reaction kinetics due to the higher temperature level on
the catalytic trays, but lower OME
1
concentrations in the distillate product. This would decrease
the OME
3-5
selectivity in the OME synthesis reactor and necessarily increase the recycle streams
and, therefore, the specifc heat demand for product purifcation. The mixture is separated into
the azeotropic mixture of 95
wt
% OME
1
and 4
.
5
wt
% MeOH in the distillate (stream 10) and a
mixture of 73 wt% FA and 27 wt%H2O in the bottom (stream 11).
5.3.2 Mass balance
A summary of the overall mass balance of the COMET process, as well as process P1 to P4 is
listed in Table 5.2 for a production capacity of 100 kt a�1 OME3-5.
The mass balance evaluation of P1-P4 shows that P1 and P3 require more H
2
and less CO
2
feedstock in comparison to P2 and P4 to produce the targeted 100
kt a�1
OME
3-5
. In fact, this
is the outcome of the two diferent process design concepts for the production of FA relying on
aqueous or monomeric FA. Considering the aqueous routes, they are characterized by a higher
production of H
2
O, which is the by-product of the acetalization reaction and exits the process as
wastewater streams in the case of P1 and P3. Moreover, P1 and P3 have smaller exhaust gas
fows due to the use of O
2
as oxidizing agent for the FA(aq.) sub-process. In contrast, in the
monomeric FA sub-process considered in P2 and P4, N
2
is used as a carrier for the feedstock
77
5 Results and Discussion
Table 5.2:
Overall mass balance for the production of OME
3-5
following the COMET process and
the processes P1 to P4. The processes were simulated with a capacity of 100
kt a�1
OME3-5 [4].
COMET P1 P2 P3 P4
Total input in kg kg�1
OME3�5 6.6 7.54 8.19 7.58 8.53
H2 0.25 0.27 0.21 0.27 0.21
CO2 1.82 1.96 2.18 1.94 2.2
Aira 4.53 5.32 5.6 5.37 5.92
Total output in kg kg�1
OME3�5 6.6 7.54 8.19 7.58 8.53
OME3-5 1 1 1 1 1
OME3 0.5 0.43 0.46 0.44 0.43
OME4 0.32 0.38 0.35 0.36 0.36
OME5 0.18 0.19 0.19 0.18 0.21
Wastewater 1.03 1.3 0.98 1.28 1
aq. FA solution 0.18 - - - -
Exhaust gas 4.39 5.24 6.21 5.3 6.54
a Air used for the FA(aq.) synthesis and for the combustion of purge streams, while the generated heat was
utilized in the processes, as shown in Figure 5.15-5.18.
MeOH, which should be introduced at certain dilution to the FA reactor. As a result of purging
a portion of the carrier gas to prevent the accumulation of the side product CO, the monomeric
FA sub-process has a higher exhaust gas fow. In addition, the side product CO of the monomeric
FA sub-process leads to a higher demand of CO
2
for P2 and P4. Alternatively, the H
2
side
product of the endothermic MeOH dissociation reaction in the monomeric FA synthesis - which
is recycled to the MeOH synthesis - lowers the demand for the total process H
2
feedstock in
comparison to P1 and P3. Consequently, this results in higher input and output mass fows for
P2 and P4. In addition, the OME
3
, OME
4
and OME
5
compositions reveal small diferences
between the process routes. However, this investigation focused on similar product compositions
rather than minimal recycle fows to defne the feedstock composition of the OME
3-5
sub-process.
This approach is based on the assumption that product composition is of greater importance to
the application than the process energy efciency of the production process.
Furthermore, the results show that the overall COMET process requires less H
2
than P1 and P3
but more H
2
than P2 and P4. The diference to P1 is mainly based on the FA concentration
sub-process, in which the simulation of the COMET process contains a modifed separation of FA
from H
2
O due to the addition of a distillation column and a third evaporator. This results in a
smaller amount of FA, which exits the process in the form of an aqueous FA solution by-product
stream (see stream 2 in Figure 4.7).
The diference to P2 and P4 is mainly based on the advantages of the anhydrous FA synthesis
from MeOH which produces H
2
as a by-product which can be separated and recycled to the
MeOH sub-process. The state-of-the-art partial oxidation of MeOH for the production of FA on
78
5 Results and Discussion
the other hand produces H
2
O. The lower CO
2
demand of the COMET process in comparison to
P1 and P4 is also based on the FA concentration sub-process and the anhydrous FA synthesis.
P1 and P3 require more CO
2
due to the higher amount of FA in the by-product stream (see
stream 2 in Figure 4.7). The lower demand of air of the COMET process is mainly a result of
the consideration of smaller purge streams which are oxidized in the combustion sub-process.
The oxygen demand for the partial oxidation of MeOH towards FA(aq.) is only slightly lower for
the COMET process than for P1 and P3.
The composition of the fnal OME
3-5
product mixture also shows signifcant diferences. While
the investigation for P1 to P4 focused on a composition close to the highest yield of OME
3-5
after the synthesis but still similar between the considered processes, the composition of the
COMET process was selected to meet the requirements for the pre-standard DIN/TS 51699.
Regarding the wastewater production, the COMET process produces less wastewater than P1
and P3 but more wastewater than P2 and P4. The diference to P1 is mainly based on the
composition of the wastewater. While the simulation of the COMET process produces high-purity
wastewater and an aqueous FA solution by-product, the simulation of P1 and P3 considered the
aqueous FA solution to be part of the wastewater. The diference to P2 and P4 is also explained
by the anhydrous FA synthesis.
The exhaust gas fow is lower for the COMET process than for P1 to P4 which is the result of
the smaller purge streams and, therefore, the lower air demand for the combustion.
5.3.3 Energy demand
The specifc energy demand and operation conditions for the main process units evaluated by
the COMET process simulation are listed in the appendix in section A.3.4.
A summary of the overall energy demand of the COMET process, as well as the processes P1 to
P4, after the heat integration is listed in Table 5.3. Besides the energy content of the feedstock
and product based on the LHV, it shows the demand for electricity, low pressure steam (LPS),
medium pressure steam (MPS), cooling water and heat above 250
◦C
in relation to the production
of 1 kg OME3-5.
The diferent H
2
demands between the COMET process and P1 to P4 directly refect on the
total process energy demand. Furthermore, the electricity demand of the COMET process is
higher than for P1 and P3 but lower than for P2 and P4. Compared to P1 to P4, the operation
conditions of the phase separators in the MeOH sub-process were adjusted resulting in higher
recycling rates and, therefore, higher compression demand. Furthermore, P1 to P4 did not
consider the compression demand for the combustion sub-process. The higher electricity demand
of P2 and P4 is a result of the anhydrous FA synthesis, which requires higher recycle streams
and dilution rates in comparison to the partial oxidation of MeOH in P1, P3 and the COMET
process and heat at high temperature level above 250 ◦C.
The demand for LPS of the COMET process is higher than the demand of P1 or P2 but lower
than the demand of P4, which is mainly a result of the heat integration strategies. P1 and P2
generate more LPS than they consume, while P3, P4 and the COMET process show higher
demands than generated. However, the demand for MPS is lower for the COMET process. P1
79
5 Results and Discussion
Table 5.3:
Overall energy demand for the production of OME
3-5
following the COMET process
and the processes P1 to P4. The processes were simulated with a capacity of 100
kt a�1
OME3-5 [4].
COMET P1 P2 P3 P4
Total input kW h kW h�1
OME3�5,LHV
H2 1.6 1.7 1.33 1.69 1.34
Total output kW h kW h�1
OME3�5,LHV
OME3-5 1 1 1 1 1
Energy demand kW h kW h�1
OM E3�5,LHV
Electricity 0.11 0.09 0.13 0.09 0.14
LPS, 4 bar 0.09 -0.1 -0.07 0.09 0.24
MPS, 23 bar 0.05 0.3 0.26 -0.16 -0.07
Cooling water -1.02 -1.05 -0.91 -1.11 -0.79
Heat T > 250 ◦C - - 0.19 - 0.19
and P2 show very high demands of MPS, while P3 and P4 generate more than they consume.
The MPS demand for the COMET process is signifcantly lower than for P1 or P2. Due to the
formation of side products, diferent feedstocks for the OME
3-5
sub-process result in diferent
recycle streams and, therefore, mass fows for the separation in the distillation columns, which
in turn leads to diferent reboiler and condenser demands. The main consumers of MPS are
CO-1 and CO-3 in the OME
3-5
sub-process. However, MPS is also generated in the MeOH
synthesis reactor and the combustion sub-process. While the combustion sub-process generates a
similar amount of MPS of about 0
.
15
kWh kWh�1
comparing P1 and the COMET
OME3�5,LHV
process, the amount difers for the MeOH reactor with
�
0
.
04 and 0
.
11
kWh kWh�1
,
OME3�5,LHV
respectively. The lower MPS generation of P1 to P4 is mainly a result of the inlet temperature to
the MeOH reactor. The inlet temperature of P1 to P4 is about 185
◦C
and, therefore, needs to be
heated up to the operation temperature of 250
◦C
using generated MPS. The inlet temperature
of the COMET process simulation is about 240
◦C
, which requires a larger heat transfer area
but improves the energy efciency. Furthermore, the demand for MPS of the distillation columns
in the OME
3-5
sub-process difer signifcantly. P1 requires about 0
.
32
kWh kWh�1
of
OME3�5,LHV
MPS for the purifcation of the OME
3-5
product stream, while the COMET process requires
only 0
.
20
kWh kWh�1
. This is mainly a result of the OME
3-5
yield after the reactor
OME3�5,LHV
as discussed in section 5.3.5.
The demand for cooling water is similar between P1, P3 and the COMET process but signifcantly
lower for P2 and P4. The cooling is required mainly for the temperature level between 90
◦C
and 30
◦C
and is, therefore, hardly utilizable for the heat integration of the COMET process.
The utilization of heat pumps instead of cooling water is evaluated in section 5.4 and shows a
signifcant enhancement potential for the overall energy efciency.
80
5 Results and Discussion
Only P2 and P4 have a demand for heat above 250
◦C
, due to the endothermic anhydrous FA
synthesis.
Regarding the heat demand for the separation of H
2
O via reactive distillation, the COMET
process requires about 1
.
1
kWh kg�1
at 117
◦C
, which is similar for the H
2
O using membranes
H2O
but lower than the H
2
O separation using adsorption, as discussed in section 2.5. The heat
demand for the separation of H
2
O via reactive distillation is based on the assumption that the
main target of the reactive distillation column is the separation of H
2
O from the loop. Therefore,
the heat demand of the reboiler and the feed preheater can be allocated to the amount of H
2
O
separated from the loop.
5.3.4 Process efciencies
A summary of the overall process efciencies of the COMET process, as well as the processes P1
to P4, after the heat integration is listed in Table 5.4.
Table 5.4:
Overall process efciencies for the production of OME
3-5
following the COMET process
and the processes P1 to P4. The processes were simulated with a capacity of 100
kt a�1
OME3-5 [4].
COMET P1 P2 P3 P4
ηenergy in %
ηC in %
ηmass in %
54.1
88.1
41.1
50.3
81.6
38.1
54.6
73.2
41.9
49.3
82.1
38.5
54.4
72.5
41.4
Process routes P2 and P4 comprising the anhydrous FA synthesis sub-process exhibit the highest
energetic efciencies due to the recycling of the valuable side product H
2
as a feedstock. In
contrast, P2 and P4 exhibit lower carbon efciencies principally due to the side reaction in the
synthesis of monomeric FA to CO, see eqn. 4.8. Evidently, as shown in Table 5.3, the lower
energetic efciencies of P1 and P3 arise principally from the higher H
2
demand, which is not
fully compensated for by the heat required at above 250
◦C
in P2 and P4. The lower overall
material efciency,
ηmass
, of the conversion of feedstock to OME
3-5
for P1 and P3 is a result
of the production of large amounts of the side product H
2
O in the synthesis of FA(aq.) due
to the MeOH partial oxidation reaction, see eqn. 4.4. This H
2
O is separated with large efort
downstream to the FA(aq.) sub-process and leaves the process in the form of a wastewater and
by-product stream.
Held et al. [109] investigated diferent scenarios to produce OME
3-5
based on stoichiometric
material balances together with diferent heat integration strategies within the sub-processes
and carbon capture scenarios for the feedstock CO
2
. Particularly, one scenario allows for heat
integration between all sub-processes in combination with CO
2
from point sources assuming
CO
2
is available without an additional energy demand. This scenario is consistent with the CO
2
feedstock assumptions from this work, which consider purchasing already prepared CO
2
without
extending the system boundaries to include the separation and preparation of CO
2
. Under this
81
5 Results and Discussion
scenario, a process energy efciency of 59
�
60% was estimated, which is higher than the process
energy efciency estimated in this work. The diference is particularly a result of the diferent
level of detail considered for the process simulation in both studies. Schemme et al. [119] and
Burre et al. [130] also investigated diferent routes to produce OME
3-5
based on H
2
feedstock.
However, in these studies, a process energy efciency of 31
�
40 % was estimated, which is
signifcantly lower than the result of this work, being closer to the results from the scenario
reported Held et al. [109], in which heat integration is only considered within the sub-processes
themselves rather than within the entire process chain. Hence, this highlights the impact and
importance of heat integration on process energy efciency with respect to the entire process
route. Specifcally, the efect of using the excess heat from the MeOH sub-process throughout
the entire process heat integration has a positive impact on the process energy efciency.
The overall energy efciency of the COMET process is higher than P1 and P3 and similar to P2
and P4. Furthermore, the carbon efciency is considerably higher than P1 to P4. The lower
carbon efciency of P1 and P3 is mainly a result of the more efcient H
2
O separation of the FA
concentration sub-process considered for the COMET process simulation. As a result, the carbon
efciency of P1 and P3 could also be increased by adjusting the FA concentration sub-process
that is to be considered in a future work. The OME
3-5
yield based on the feedstock H
2
and
CO
2
is also higher for the COMET process than for P1 and P3 which is also a result of the
more efcient H
2
O separation of the FA concentration sub-process. The OME
3-5
yield is similar
to P2 and P4 since H
2
O, formed in the FA(aq.) synthesis, is separated from the process loop,
compared to the formation, separation and recycling of H
2
in the anhydrous FA synthesis of P4.
5.3.5 Comparison of alternative OME3-5 production processes
To compare alternative OME
3-5
production process concepts, important performance parameters
are listed in Table 5.5. These performance parameters include the mass fraction of OME
3-5
before and after the OME reactor
wOME3�5
. Furthermore, the heat demand of the reboiler of
the distillation columns and feed preheaters of the OME sub-process in relation to the OME
3-5
product massfow times its LHV
QReboiler/HOME3�5
is considered. Another key performance
parameter is the overall energy efciency
ηenergy,overall
. This considers the entire process chain
starting from H
2
O electrolysis and CO
2
via the production of the intermediate products towards
the target product mixture OME
3-5
. Regarding the electricity and heat demand for the H
2
O
electrolysis and CO
2
preparation, the assumptions from Held et al. [109] were considered. For
the CO
2
preparation all three scenarios from Held et al. [109] were considered, comprising CO
2
from point sources (CPS), post combustion capture (PCC) using mono-ethanol amine scrubbing
and direct air capture. The key assumptions for the expanded system boundary evaluation are
summarized in the appendix in section A.3.5. In addition, the scale-up potential in the near
future is evaluated. The key performance parameters are based on the results of this work, Held
et al. [109] and Schemme et al. [119].
Regarding the yield of OME
3-5
after the reactor as illustrated in Figure 4.1, the anhydrous process
concepts show far higher OME
3-5
concentrations than the aqueous process concepts, as indicated
by
wOME3�5
. This also refects on the heat demand for the OME
3-5
product purifcation, which
82
5 Results and Discussion
Table 5.5:
Comparison of various OME
3-5
production processes based on the results of this work
[4], Held et al. [109] and Schemme et al. [119].
Anhydrous synthesis Aqueous synthesis
Feedstock
OME
1
DME OME
1
and MeOH and OME
1
and MeOH and
COMET
and TRI and TRI monomeric FA(aq.) FA(aq.) monomeric
FA or pFA FA
wOM E3�5 in wt%
5-34 [109]
0-35 [119]
-
a
5-29
QReboiler
/
HOME3�5 in kWHeat
kW �1
OM E3�5
ηenergy,overall in %
Scale-up potential in the near
future
7
.
6 % [109]
5
.
5 % [119]
29-37 [109]
22-26 [119]
likely
-
a
-
a
unlikely
15 %
27-36
unlikely
0-15 [109]
4-16
4-19 3-19 0-20
0-14 [119]
47 % [109]
39 %
26 % 48 % 35 %
78 % [119]
30-36 [109]
25-31
26-32 28-37 28-34
less likely less likely unlikely likely
a Further investigations and an adjusted process concept are required to estimate the process performance.
is compared based on
QReboiler/HOME3�5
. The anhydrous process concepts show signifcantly
lower heat demands for the product purifcation. Exceptions are the production based on OME
1
and FA(aq.) or pFA and the COMET process, which despite comparatively low yields of OME
3-5
in the reactor require less heat for the separation of the target product than the other aqueous
process concepts. For a consistent basis of comparison, the production of the intermediate
products for diferent OME production processes was included into the evaluation, indicated by
ηenergy,overall
. The result is a low overall energy efciency of
<
40 % for all processes, with minor
diferences between anhydrous and aqueous process concepts. Greater diferences were reported
between diferent literature sources, which is especially signifcant comparing the results for the
OME
3-5
production based on OME
1
and TRI as well as MeOH and FA(aq.) from this work,
Held et al. [109] and Schemme et al. [119]. Those diferences are discussed in detail in section
5.3.1 and mainly result from diferent heat integration strategies and the simulation procedure.
Schemme et al. [119] only integrated the heat between individual sub-processes, the simulation
in this work and by Held et al. [109] considered the heat integration between all sub-processes.
However, Held et al. [109] did use stoichiometric material balances and literature data, while
the processes in this work and by Schemme et al. [119] were simulated with the software Aspen
Plus®.
Regarding the low overall energy efciency of all OME
3-5
production processes, the following
section shows the potential of including HTHP to lift the temperature of the excess heat streams
and, therefore, supply internal heat demands and, in addition, external heat demands. Besides
only small diferences in the energy efciency, the production costs of the OME
3-5
product also
show no signifcant diferences between diferent production processes [1, 119].
Regarding a sustainable large-scale production of OME
3-5
in the near future only the COMET
process and a production based on OME
1
and TRI can already be scaled up today. However,
the latter is comparatively complex, comprising fve sub-processes for the production of MeOH,
FA(aq.), TRI, OME
1
and OME
3-5
. A sustainable OME
3-5
production based on the COMET
process on the other hand comprises three sub-processes for the production of MeOH, FA(aq.) and
OME
3-5
. The OME production process based on DME and TRI requires further investigations,
mainly due to the high MEFO formation during the synthesis and the low activities for the
83
5 Results and Discussion
conversion of DME to OME. A fast scale-up of the processes based on monomeric FA is mainly
prevented by the low TRL of the monomeric FA production. Finally, the aqueous process
concepts require the separation of H
2
O from the loop of the OME
3-5
sub-process, which is the
main bottleneck for a fast scale-up. Various concepts for separating H
2
O from the loop were
already proposed, and some show promising results, as discussed in section 2.5 and demonstrated
in this work for the reactive distillation column. This enables a scale-up for the processes based
on MeOH and FA(aq.) and OME
1
and FA(aq.) or pFA. In comparison to the OME
3-5
production
based on OME
1
and TRI, the aqueous process concepts enable a considerable simplifcation,
which typically improves the robustness and, therefore, feasibility for large-scale application.
84
5 Results and Discussion
5.4 Process evaluation for an improved energy efciency of Power-to-X
processes using heat pumps
The following sections present the results form [2].
The extended system boundary for the evaluation of the improved energy efciency of the
production of OME3-5 is illustrated in Figure 3.6 in section 3.4.
Renewable electricity is the basis for the PtX value chain. It supplies the power demand for
all sub-processes, including the production of OME
3-5
from H
2
and CO
2
, the preparation of
H
2
via H
2
O electrolysis, the seawater preparation, the capture of CO
2
from ambient air, and
the process heat provision. In addition, heat streams are exchanged between the sub-processes,
which increases the overall process energy efciency.
Seawater is desalinated, purifed, and electrochemically split into its elements O
2
and H
2
via H
2
O
electrolysis. This provides the feedstock H
2
. In parallel, the feedstock CO
2
is supplied by fltering
it from the air using DAC technologies. The CO
2
, which is thermodynamically very stable, is
activated with the energy-rich H
2
in a thermochemical catalytic process at elevated pressure and
temperature levels to produce MeOH. One part of the MeOH is further converted to FA, and
the remaining MeOH reacts with FA to produce OME. Finally, from the OME synthesis product
mixture, the target product OME
3-5
is purifed in a cascade of thermal separation units. Heat
streams are integrated between the sub-processes, but a signifcant amount of low-temperature
excess heat still leaves the process. To enable a suitable application for this low-temperature
heat, HTHP are introduced. The HTHP are supplied with renewable electricity and lift the
temperature level of the excess heat streams above 100
◦C
. This enables a higher share of heat
integration between the sub-processes and additionally supplies further external applications.
Some suitable applications were identifed and evaluated in terms of temperature level and heat
demand. Therefore, besides the production of the target PtX product, excess heat streams are
produced as additional product streams and were considered in the overall process evaluation.
In the following, the OME value chain will be introduced stepwise, starting from the OME
production from H
2
and CO
2
as primarily evaluated in literature and further expanding the
system boundaries to include the main conversion steps starting from the primary energy and
raw materials seawater, air, and renewable electricity.
5.4.1 OME3-5 production from H2 and CO2 (system boundary I)
System boundary I considers the production of OME
3-5
from H
2
and CO
2
and is based on the
results from section 5.3.2, 5.3.3 and 5.3.4 [1].
A simplifed process fow diagram is illustrated in Figure 5.15, which is accompanied by a
detailed description of the OME
3-5
production process based on H
2
and CO
2
via the intermediate
production of MeOH and FA(aq.), see section 5.3.1.1.
In section 5.3.4 the overall energy efciency for the OME
3-5
production (scenario I) based on
MeOH and FA(aq.) was evaluated with 50 %, following eqn. 3.6 and, therefore, not considering
the excess steam at 4
bar
as a by-product of the production process. Following eqn. 3.9,
85
5 Results and Discussion
the overall energy efciency of the process increases to 53 %. Table 5.6 summarizes incoming
energy streams, educt and product streams, and the energy efciency following eqn. 3.9 for the
production of OME
3-5
based on the results of section 5.3.2, 5.3.3 and 5.3.4 [1]. The OME
3-5
product stream and the feed streams are presented as the product of mass fow and lower heating
value (LHV). The LHV of OME
3-5
and H
2
were taken from Held et al. [109] with 5
.
25
kWh kg�1
OME3-5 and 33.33 kWh kg�1 H2, respectively.
Table 5.6:
Key performance indicators of the process for the production of 100
kt a�1
OME
3-5
from H2 and CO2 via MeOH and FA(aq.) [1].
Energy demand and content Mass fow rate
[MW ] [kW h kW h�1
OME3�5 ] [t h�1] [t t�1
OME3�5 ]
Products
OME3-5 65.6 1 12.5 1
Steam, 4 bar 6.7 0.1
Feed
H2 111.4 1.7 3.3 0.27
CO2 0 0 24.4 1.96
Energy demand
Electricity 5.9 0.09
Steam, 20 bar 19.9 0.3
Energy efciency, eqn. 3.9 [%] 52.7
The energy efciency of the production process of OME
3-5
changes by considering diferent system
boundaries, as presented in Figure 3.6 and can be signifcantly increased by recovering excess
heat streams using HTHP. Figure 5.19 illustrates the diference in the energy efciency between
the diferent system boundaries and shows that the overall energy efciency can be lifted to 61 %
and above, if the recovered excess heat can be used as a product heat stream to supply the heat
demand of other processes or process steps. This is described in detail in the following sections.
5.4.2 Including H2 production (system boundary II)
The system expansion to consider the H
2
production using H
2
O electrolysis, as illustrated in
Figure 3.6, signifcantly infuences the process energy efciency. Today, H
2
is mainly produced
via steam methane reforming (SMR) [154]. However, using renewable electricity, H
2
can be
produced with low carbon footprints using H
2
O electrolysis technologies [155]. The leading H
2
O
electrolysis technologies with already commercially available systems on MW scale are polymer
electrolyte membrane (PEM) electrolysis, alkaline electrolysis (AEL), and solid oxide electrolysis
cells (SOEC). Some examples of commercially available electrolysis systems are presented in
Table 5.7. In comparison to the demand of 3
.
3
t h�1
H
2
for a production of 100
kt a�1
OME
3-5
,
an electrolyzer capacity of around 174
MW
will be needed at full load hours at the assumed
efciency, see Table 5.8. However, as shown in Table 5.7, the sizes of the already available H
2
O
electrolysis systems require further scale-up toward large-scale PtX production processes.
86
5 Results and Discussion
Figure 5.19:
Process energy efciency progress for producing OME
3-5
from H
2
and CO
2
including
H
2
O electrolysis, seawater desalination, and CO
2
capture directly from air. The
process is heat-integrated, and low-temperature product streams are lifted to 100
◦C
using HTHP and supply the heat demand for DAC, seawater desalination, and other
applications [2].
While PEM and AEL systems operate at low temperatures
<
90
◦C
, SOEC systems operate
at very high temperatures of 700
�
900
◦C
. The latter technology is more attractive at high
temperature heat availability, which is not the case for the considered OME value chain system
boundary. That is why PEM and AEL systems were considered in this evaluation for the
production of H
2
. Table 5.8 presents the key operational parameters for the H
2
O electrolysis
considered in this work.
With an efciency of 64 %, the specifc electricity consumption for the production of H
2
is
52
.
1
MWhel t�1
. Regarding the useful excess heat from the electrolysis, almost no information
H2
is available for PEM and AEL H
2
O electrolysis systems. Tiktak [161] simulated PEM stacks
and estimated that 18
.
4 % of the power consumption by the stack could be extracted as usable
excess heat at 80
◦C
. Holst [162] simulated an AEL system at 70
◦C
and estimated that 15
.
8 %
of the power consumption could be extracted as usable excess heat. Zhang et al. [163] conducted
a life cycle assessment of Power-to-Gas systems and assumed a waste heat from the electrolysis
of about 28
.
5 %, which they used for heat integration. Bergins et al. [164] investigated the
potential of upgrading waste heat from electrolysis and a post combustion CO
2
capture plant
using HTHP to produce MeOH. They assumed that 15
.
9 % of the power consumption could be
extracted as usable excess heat at 80
◦C
. A report evaluating the synergies between PtX and
district heating (DH) in Denmark assumes that 25 % of the power consumption of the electrolysis
87
5 Results and Discussion
Table 5.7:
Some commercially available H
2
O electrolysis systems and their energy and H
2
O
demand [2].
Type
System, Company
System
Production
power rate
[
MW
]
[
kgH2 h�1
]
Efciency Specifc Specifc
Operation References
(LHV H
2
electricity H
2
O con- temper-
to power demand sumption ature
�1
input) [%]
[
MW hel t
]
kgH2O kg�1
[
◦C
]
H2 H2
AEL
Sunfre-Hylink Alka-
line, Sunfre GmbH
AEL
thyssenkrupp
Uhde Chlorine
Engineers GmbH
PEM
Silyzer 300, Siemens
Energy AG
PEM
HyLYZER
®
-4000,
Cummins Inc.
SOEC
Sunfre-Hylink SOEC,
Sunfre GmbH
10
20
17.5
18
2.7
200
359
335
359
67
64
64
64
65
84
b)
52
50
52
51
40
9.5
a) ≤
85 [156]
< 11
a) ≤
90 [157]
10
a)
50-80 [154,
158]
9
a)
50-80 [154,
159]
12.8
b)
700-900 [154,
160]
a) DI H2O
b) steam at 150 � 200 ◦C is used as H2O source, which is not included in the energy efciency evaluation.
Table 5.8:
Key operation parameters of the H
2
production via H
2
O electrolysis considered in this
work [2].
Technology PEM/AEL
Efciency (LHV H2 to power input) [%] 64
Specifc electricity consumption MW hel t�1 52.1
H2
Usable excess heat [%] 20
Specifc usable excess heat MW hth t�1 10.4
H2
Temperature of excess heat [◦C] 70
Specifc H2O consumption kgH2O kg�1 10
H2
system, including compression, is produced as heat at 35
�
70
◦C
[165]. In this work, 20 % of the
power input of the electrolysis system is assumed to be extracted as usable excess heat streams
at 70
◦C
, which leads to 10
.
4
kW hth
per kg H
2
produced. For the efciency estimation of this
system boundary II, only the electricity demand for the production of H
2
was included. The
usable excess heat will be considered in system boundary IV after a temperature lift using HTHP.
Including the H
2
production of 0
.
27
kgH2 kg�1
, the process energy efciency decreases by
OME3�5
17 % to reach an overall 36 %. Held et al. [109] also investigated the energy efciency of the
production of OME
3-5
from diferent routes considering diferent scenarios for heat integration
and CO
2
capture. For their scenario 3, route A uses CO
2
from point sources similar to system
boundary II in this work. An energy efciency of 36 % was evaluated, which agrees with the
results of this work. This scenario includes the production of OME
3-5
from H
2
and CO
2
based
on MeOH and FA(aq.) with an extensive heat integration and considers the H
2
production using
a PEM system, but without considering the energy demand for the CO2 preparation.
5.4.3 Including CO2 preparation (system boundary III)
CO
2
can be captured from point sources in industrial processes like metallurgical plants and
steelworks [166], waste incineration plants [167], power plants, refneries, steam crackers, steam
reforming processes, NH
3
production plants, cement production plants, biogas plants, and
88
5 Results and Discussion
others [168]. The capacity of these point sources in Europe today is about 1477
Mt
CO
2
.
Considering the transformation towards a low carbon economy, the capacity can reduce to about
330
Mt
CO
2
[169]. Alternatively, CO
2
can be captured from seawater or air, enabling the use
of location-independent and long-term sources and an almost closed carbon loop. Despite an
efective concentration of 0
.
099
kgCO2 m�3
in seawater compared to the concentration of about
0
.
00079
kgCO2 m�3
in air, currently, the technology for capturing CO
2
from seawater is not yet
mature. On the other hand, there are several demonstration plants for direct air capture of
CO
2
[170, 171]. The leading DAC technologies with already commercially available systems are
based on high temperature (HT) aqueous solutions with KOH or temperature-vacuum swing
adsorption (TSA). Some examples of commercially available DAC systems are presented in Table
5.9, showing that demonstration plants are already in operation capturing up to 4
ktCO2 a�1
but
also indicating that for capturing 24
.
4
tCO2 h�1
for a production of 100
ktOME3�5 a�1
, further
scale-up is required for large-scale PtX production processes. Despite a signifcantly higher
energy demand for capturing CO
2
from air instead of using point sources, today, DAC can reach
carbon footprints of
< �
0
.
93
kgCO2,eq
per kg CO
2
captured using low-carbon energy sources [49].
Currently, there are 18 DAC plants in operation with a total capacity of
>
0
.
01
MtCO2 a�1
. By
2030, the capacity is expected to grow > 40 MtCO2 a�1 [171].
Table 5.9: Some commercially available DAC systems and their specifc energy demands [2].
Type
Company Specifc elec- Specifc heat Temperature Production References
tricity de- demand in level of the sup- rate demon-
mand in
kW hth kg�1
plied heat in
◦
C
stration plants
�1 CO2
kW hel kg h�1
CO2
in
tCO2
HT aqueous
solution
Carbon Engi-
neering Ltd.
0.37 1.46 900 0.05 [172–176]
TSA
Climeworks AG 0.5 1.5 <100 0.5 [177, 178]
TSA
Global
Thermostat
0.26 1.4 85-95 0.09-0.5 [173, 175,
176, 179]
While TSA systems operate at low temperatures
<
100
◦C
, HT aqueous solution-based systems
operate at very high temperatures of approx. 900
◦C
, similar to SOEC systems. Due to a more
comprehensive heat integration using low-temperature technologies, the CO
2
for the production of
OME
3-5
was assumed to be captured using the specifcations from Climeworks AG as presented in
Table 5.9. Excess heat from the PtX process can be used to provide the required low temperature
heat. Furthermore, the slight temperature diference of 5
�
50
◦C
(depending on the system
confguration) between the heat needed for the desorption in the TSA process for DAC systems
and the excess heat from PEM and AEL systems can be compensated using HTHP technologies,
which will be considered in the system boundary IV.
With an electricity demand of 0
.
5
kW hel kg�1
and 1
.
5
kW hth kg�1
heat demand at 100
◦C
,
CO2 CO2
the energy efciency for system boundary II drops by 7 % to reach an overall 29 %, considering
the preparation of 1
.
96
kgCO2 kg�1
. Held et al. [109] also investigated the infuence of the
OME3�5
DAC system on the energy demand of the OME
3-5
production process and evaluated an overall
process energy efciency of approx. 30 %.
89
5 Results and Discussion
5.4.4 Including HTHP systems (system boundary IV)
HTHP or very high temperature heat pumps (VHTHP) with sink temperatures above 80
◦C
show considerable potential for excess heat recovery to supply heat to various industrial processes
such as steam generation, food preparation, and distillation processes [180, 181]. In Table 5.10,
examples of commercially available HTHP systems are listed with information about heating
capacity, sink temperature, temperature lift, and the corresponding coefcient of performance
(COP).
Table 5.10: Some commercially available HTHP systems and their performance [2].
Type
Company Max. heat- Sink tem-
Temperature
COP
References
ing capac- perature lift in
K
ity in
MW
in
◦C
Hybrid heat pump Hybrid En- 2.5 100 60 4.5 [180]
(absorption and ergy AS
compression)
Compression
Kobelco 0.66 120 55 3.5 [180]
Compressors
Co., Ltd.
Compression
Heaten AS 1 150 50 4.6 [180, 182]
Compression
MAN Energy 48.4 110 70 3 [183]
Solutions
The HTHP systems in Table 5.10 show that excess heat can be recovered down to 40
◦C
and
below with a considerably high COP of up to 4.5. The resulting sink temperature of 100
◦C
and above can directly be used for various applications or even raised to 150
◦C
using a second
HTHP system. For the evaluations in this work, the performance of the hybrid heat pump
system from Hybrid Energy AS with source temperatures down to 40
◦C
and a constant COP
of 4.5 is considered. Regarding the recovery of the excess heat from the H
2
O electrolysis using
HTHP systems, the heat demand of 37
MW
for the DAC system is already exceeded, see Table
5.11. Only covering the heat demand from the DAC system leads to an energy efciency of
33 % (system boundary IVa) and, therefore, a lift of 4 % in comparison to system boundary
III. Providing the rest of the excess heat from the H
2
O electrolysis as a product heat stream at
100
◦C
lifts the energy efciency by another 3 % to reach an overall process energy efciency of
36 % for system boundary IVb.
Another promising excess heat source is the cooling demand for the condensation in the distillation
columns. Recovering this condensation enthalpy leads to an additional product heat stream of
51
MW
at 100
◦C
, which lifts the overall energy efciency to 56 % for system boundary IVc.
Furthermore, using the cooling demand before the frst fash unit in the MeOH sub-process,
the cooling demand for the condensation within the FA concentration sub-process and reducing
the temperature of the exhaust gas stream leads to another additional product heat stream of
14 MW at 100 ◦C and lifts the energy efciency to 61 % for system boundary IVd.
The OME
3-5
production process contains further streams which need to be cooled and could
potentially supply HTHP systems. However, these cooling demands are either comparatively
90
5 Results and Discussion
Table 5.11:
Waste heat recovery using HTHP systems and their infuence on the overall process
energy efciency, assuming a constant COP of 4
.
5 for temperature lifts from down to
40 to 100 ◦C [2].
Excess heat source
H
2
O elec- Condenser
Cooling warm
trolysis (IVa
distillation
streams (IVd)
and IVb) columns (IVc)
Excess heat source in MWth
35 40 11
Additional electricity demand in MWel
10 11 3
Recovered heat in MWth
45 51 14
Overall energy efciency for the produc-
36.2 56.3 61.4
tion of OME3-5 in %
small such as the intercooling within the feed stream compression or the cooling demand of
the OME reactor inside the OME
3-5
sub-process, or they require very low temperatures below
40
◦C
, such as the cooling demand for the FA absorption inside the FA sub-process. In total,
the recovery of suitable excess heat streams along the production process of OME
3-5
leads to
additional product heat streams of 110
MW
at 100
◦C
or 73
MW
if only the heat demand of
DAC is covered. However, these heat streams can only be accounted for as product streams if
they can be used for other processes or process steps. This is considered in process boundaries
Va and Vb.
5.4.5 Potential applications for the recovered excess heat (system boundary V)
Besides the utilization of a proportion of the recovered excess heat for the DAC system, the heat
could be used for other applications such as seawater desalination, DH, and steam generation,
depending on the demand close to the production site.
Seawater desalination is mainly based on reverse osmosis (RO), which only requires electrical
energy and no heat and accounts for 68
.
7 % of the globally installed capacity in 2019. Multi-stage
fash desalination (MSF) and multi-efect distillation (MED), on the other hand, require low-
temperature heat at around 100
◦C
and still account for 17
.
6 % and 6
.
9 % of the globally installed
capacity, respectively [184]. Both technologies show a similar heat demand of 30
�
120
kWh m�3
H2O
but diferent electricity demands of 1
.
5
�
2
.
5
kWh m�3
for MED and 3
�
6
kWh m�3
for
H2O H2O
MSF in comparison to 2
.
5
�
7
kWh m�3
for RO [184–186]. To produce 1
kgH2
, about 10
kg
H2O
of H
2
O is required, which leads to a consumption of approx. 2
.
7
kgH2O kg�1
for the H
2
O
OME3�5
electrolysis. Assuming that the H
2
O is prepared using MED with an average heat demand of
50 kWh m�3 , only 1.5 % of the product heat stream of 110 MW at 100 ◦C is required.
H2O
District heating systems can be used to defossilize the heat demand of the building sector and,
therefore, reduce GHG emissions if renewable or excess heat sources are used. Besides industrial
excess heat, large-scale heat pumps can valorize process excess heat to cover this sectors heat
demand. In Stockholm, heat pumps have been supplying the DH network since the 1970s, with a
total heating capacity of 660
MWth
today. With temperatures up to 115
◦C
, the DH network of
91
5 Results and Discussion
Stockholm has a dimensioning load of 4
.
8
GWth
[187]. The greater Copenhagen area DH network
has a peak load of 2
.
5
GWth
with temperatures up to 100
◦C
[188]. Furthermore, large-scale heat
pumps are estimated to supply 25
�
30 % of the total DH heat demand in Europe by 2050, which
accounts for approx. 40
GWth
[189, 190]. In 2017, about 1
.
6
GWth
were supplied by large-scale
heat pumps with capacities between 3
�
19
MWth
[190]. On the one hand, this shows that the
capacity and temperature level of the low-temperature product heat stream from the OME
3-5
production process is suitable for DH networks. On the other hand, the required capacity of
heat pump systems to provide the product heat streams is state-of-the-art and has already been
in operation for several decades.
Using a second heat pump such as the HeatBooster from Heaten AS, the product heat stream at
100 ◦C can produce steam at 150 ◦C, which can be used for various industrial processes.
Figure 5.20 shows that the low-temperature product heat streams of the OME
3-5
production
process can be used to supply the heat demand for the DAC system and the seawater de-
salination, which only covers about 35 % of the available heat capacity. Therefore, other
applications such as DH networks or steam generation should be supplied to use the full potential
of the product heat stream and achieve energy efciencies of
>
61 % for the production of OME
3-5
.
Figure 5.20:
Utilization of the low temperature product heat stream from the OME
3-5
production
process and H2O electrolysis upgraded using HTHP [2].
5.4.6 Potential analysis: Upgrading the excess heat of the H2O electrolysis using HTHP to
supply the low-temperature heat demand of DAC systems for various PtX products
Considering the supply of usable excess heat from the H
2
O electrolysis of 10
.
4
kWh kg�1
H2
(20 % of the power demand) at 70
◦C
to an HTHP with a COP of 4
.
5 to be lifted to 100
◦C
,
13.4 kWh kg�1 can be provided at 100 ◦C. With a demand of 1.5 kWh kg�1 heat for a DAC
H2 CO2
system, 8
.
9
kgCO2 kg�1
can be provided only using the recovered and lifted excess heat from the
H2
92
5 Results and Discussion
H
2
O electrolysis. In Table 5.12, the ratio of the demand of CO
2
to H
2
is presented for various
PtX products, based on published results from process simulations [1, 119]. It shows that for all
considered PtX products, the heat demand of the DAC system can be supplied by the recovered
and lifted excess heat of a PEM or AEL electrolyzer.
Table 5.12:
H
2
and CO
2
demand for various PtX products and the recovery of the excess heat
from a PEM or AEL electrolyzer with 10
.
4
kWh kg�1
using HTHP with a COP of
H2
4.5 to supply the heat demand of the DAC system of 1.5 kWh kg�1 [1, 2, 119].
CO2
PtX product
DME
FT-diesel
fuel
Methane MeOH
MtG
OME
1
OME
3-5
H2 demand in kgH2 kg�1
P tXproduct
CO
2
demand in
kgCO2 kg�1
P tXproduct
Ratio in kgCO2 kg�1
H2
Heat demand of DAC over excess
0.26
1.91
7.27
81
0.48
3.06
6.37
71
0.5
2.75
5.5
62
0.19
1.37
7.26
81
0.4
2.87
7.13
80
0.27
1.97
7.29
82
0.27
1.96
7.31
82
heat of H2O electrolysis in %
Schemme et al. [119] evaluated an energy efciency for the production of OME
3-5
from H
2
and
CO
2
of 30
.
5 %, which is lower than the efciency of 50 % evaluated in this work or the efciency
of about 60 % evaluated by Held et al. [109]. This disagreement was discussed in section 5.3.1
and is mainly based on the level of heat integration between all sub-processes or only inside
the individual sub-processes. Moreover, for the other PtX products presented in Table 5.12,
Schemme et al. [119] evaluated an overall energy efciency between 44
.
8 % and 60 %. Following
the approach presented in this work to upgrade the excess heat streams to valuable by-product
streams using HTHP, these PtX processes also show signifcant potential to achieve much higher
energy efciency than the OME value chain presented in this work.
93
6 Conclusion and outlook
The investigations of this work focus on the synthesis and production of OME
3-5
from various
feedstocks. The main outcomes are the introduction and experimental demonstration of the novel
COMET process concept for the production of OME
3-5
, which is based on the simple feedstock
MeOH and FA(aq.) and solves the challenging H
2
O separation from the loop using a state-
of-the-art reactive distillation column. Furthermore, various catalyst systems are investigated
for the OME synthesis, showing that the IER A15 and A46 are especially suitable with a low
selectivity towards side products and comparatively fast reaction kinetics. However, the results
also indicate, that a neutralization should be considered after the OME synthesis to enable the
product purifcation using distillation columns. Moreover, various OME
3-5
production processes,
including the COMET process, are simulated and compared with literature results, showing
that starting form the sustainable production of MeOH via H
2
and CO
2
, the overall process
energy efciency for all processes is 22
�
37 %, depending on the preparation of captured CO
2
.
Thereby, a big share of the excess heat is available at low temperature levels between 30
�
90
◦C
,
which is, therefore, hardly integrable into the process heat integration. The addition of HTHP to
upgrade low-temperature excess heat streams to temperature levels higher than 100
◦C
shows
the potential to signifcantly improve the overall energy efciency to 61 % and higher, if suitable
applications are available near by the production site.
OME synthesis
For the industrial production of OME, solid acid commercially available catalysts were investi-
gated. To compare the feasibility of diferent IER, zeolites and Nafon catalysts for the OME
synthesis in an aqueous reaction system and an anhydrous reaction system, the criteria conver-
sion, selectivity, yield, activity, side product formation and thermal stability of the synthesis
product were investigated. The results show that all investigated catalysts are suitable for
the OME synthesis for both the anhydrous and the aqueous reaction systems. However, the
IER showed signifcantly higher activities and lower MEFO side product formations than the
zeolites. A15 and A36 show higher TRI side product formations, but TRI also reacts to OME
and, therefore, does not need to be separated from the loop inside the OME
3-5
production
process. Regarding the formation of FA and MeOH for the OME
1
-TRI feed mixture, all cat-
alysts led to comparatively high concentrations, which can only be prevented using very dry
feedstock. Without the separation of traces of H
2
O and MeOH inside the feed mixture of
the anhydrous reaction system, H
2
O needs to be separated from the loop to circumvent its
accumulation and negative efects on the product selectivity and reaction kinetic. Regarding
thermal stability, all catalysts indicated at least minor changes in the product composition
after the distillation and, therefore, require a neutralization step before entering the separation
cascade. In conclusion, the IER catalysts are identifed as most suitable for the OME synthesis
for both anhydrous and aqueous reaction systems, with a particular recommendation for Dowex,
95
6 Conclusion and outlook
A15 and A46 for anhydrous reaction systems and A15, A36 and A46 for aqueous reaction systems.
COMET process concept
The COMET process, which solves the challenging H
2
O management of aqueous OME processes,
was introduced in this work. The process benefts from a simple feedstock preparation, a short
process chain from H
2
and CO
2
to OME
3-5
, a comparatively high OME
3-5
yield after the reactor,
and the possibility of extracting the by-product H
2
O from the loop using a state-of-the-art
reactive distillation unit.
Other H
2
O separation methods, as discussed in the literature, were presented, and their main
advantages and hurdles were evaluated quantitatively. The main advantage of the H
2
O separation
in the COMET process via reactive distillation is the scale-up potential and the feasible application
in large-scale production plants.
Starting solely from MeOH and FA(aq.) commercial feedstocks, the main COMET process
units, comprising all evaporation, reaction and separation process steps, were experimentally
demonstrated on a pilot scale. Importantly, the technical feasibility of the reactive distillation
column - the heart of the COMET process concept - was demonstrated for a long duration of
around 600
h
on stream. In addition, the purifcation of the fnal OME
3-5
product was successfully
realized with a product compliant with the pre-standard DIN/TS 51699.
The COMET process was simulated and evaluated using Aspen Plus
®
and compared with
relevant alternative OME
3-5
production processes. Therefore, the system boundary was expanded,
including H
2
production via H
2
O electrolysis, CO
2
capture and all intermediate production
sub-processes. With an overall energy efciency of 28
�
34 %, depending on the CO
2
source, the
energy demand of the COMET process is similar to the alternative OME
3-5
production processes,
in which overall energy efciencies were evaluated in the range of 25
�
36 %. Moreover, the
COMET process shows a higher carbon efciency of 88 %.
The OME market is limited by the lack of technologically feasible large-scale processes. However,
compared to relevant alternative OME
3-5
production processes, the novel COMET process shows
the smallest technological hurdles and can already be demonstrated and scaled up.
OME3-5 production processes
Based on a standardized and validated modelling and simulation methodology implemented in
Aspen Plus
®
, diferent process routes for the production of OME
3-5
were evaluated. The feedstock
for all routes are CO
2
and H
2
to enable the sustainable production of the frst intermediate
product MeOH. Diferent sub-processes follow to prepare the intermediates for the synthesis
of OME, i.e. FA(aq.) with MeOH for P1, monomeric FA with MeOH for P2, FA(aq.) with
OME
1
for P3 and monomeric FA with OME
1
for P4. Subsequently, OME
3-5
are synthesized and
purifed on the scale of 100 kt a�1 OME3-5.
All processes were heat integrated along the process chain. Based on the mass balance and
energy demand, the overall energy efciency varied between 50
�
55 %. Processes P2 and P4
based on anhydrous FA showed a trend of higher energy efciency due to the reutilization of the
H
2
by-product from the MeOH endothermic dissociation reaction in the MeOH synthesis loop.
An important lever to enhance the energy efciency of the processes is to reduce the heat losses
96
6 Conclusion and outlook
by utilizing low-temperature excess heat through usage in external processes. A big share of the
excess heat (
>
43 %) is still at useful temperature levels considering the usage of heat pump
technology, an approach that was investigated in the frame of this work. This strategy is crucial
in the context of PtX processes, where the production will probably take place where cheap
renewable electricity is abundant. This will reduce the supply of external utility streams and
besides the overall process efciency enhacement, this can refect positively on the production
costs.
Improving the overall energy efciency by applying HTHP
The evaluation of the process energy efciency of the OME value chain considering diferent
system boundaries was presented. Starting from the system boundary considering the production
of OME
3-5
from H
2
and CO
2
and considering the excess steam as a valuable by-product, an
energy efciency of about 53 % was evaluated. Expanding the system boundary to account for
the production of H
2
via H
2
O electrolysis and the provision of CO
2
using DAC technologies,
the overall efciency drops signifcantly to about 29 %. However, most of the energy leaves the
process in form of low temperature heat streams, which can be lifted to higher, useful temperature
levels using HTHP. This provision turning a waste stream into a useful process by-product,
results in an overall process energy efciency lift to reach up to 61 %.
Besides the excess heat from the production process of the PtX product, the high capacities
of excess heat from the H
2
O electrolysis can signifcantly increase the process energy efciency
when the temperature is lifted using HTHP to meet the demand of suitable applications. This is
de facto only valid for scenarios where low temperature heat demanding processes are in the
vicinity of the PtX processes. At a temperature of about 100
◦C
, this upgraded heat streams
can be suitable for various applications such as the DAC system, seawater desalination, district
heating, or steam generation using additional HTHP. To use the full potential of the excess heat
in multiple locations, further suitable applications should be identifed to replace fossil-based
heat supplies. Besides the efciency and the potentially positive ecologic impact, the economic
aspect should be investigated to identify the trade-of between CAPEX and OPEX considering
HTHP in the value chain. Importantly, the presented examples of commercially available systems
pointed out that further scale-up is required to demonstrate the technical feasibility and reduce
the investment costs to enable a realistic economic evaluation. Furthermore, the interaction of
the PtX production, HTHP systems, and excess heat applications should be investigated and
demonstrated to identify technologically suitable and economically feasible confgurations.
97
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Acknowledgements
I would like to thank all those who have supported me professionally as well as personally on
this journey.
This thesis is the result of my work as doctoral researcher at the Institute of Chemical and
Process Engineering at TU Berlin, a reaseach scientist in the group of Dr.-Ing. Ouda Salem and
Max Hadrich at the Fraunhofer Institute for Solar Energy Systems ISE from 2018 to 2023 and
my time as Deutsche Bundesstiftung Umwelt PhD scholar (20018/541).
A special thanks goes to my dear colleagues at Fraunhofer ISE who invested plenty of their
time, energy and dedication into this successful journey. I am grateful for my supervisors Prof.
Dr. Christopher Hebling, Dr. Achim Schaadt, Dr.-Ing. Ouda Salem and Max Hadrich for their
strong support, trust, encouragement and for the freedom they gave me to develop my research.
Turning scientifc ideas into reality requires a strong, dedicated and motivated team. Therefore,
a special thanks goes to my colleagues, students and friends in the whole division Hydrogen
Technologies for their support, ideas, discussions, feedback, encouragement and humour.
I would like to thank my supervisor Prof. Dr.-Ing. Matthias Kraume at the Institute of Chemical
and Process Engineering for his strong scientifc and personal support as well as his time and
positive feedback during this journey.
Prof. Dr.-Ing. Erik von Harbou from TU Kaiserslautern is sincerely thanked for his agreement
to evaluate this thesis as second reviewer.
The acceptance for my scholarship provided by Deutsche Bundesstiftung Umwelt was a key
moment for this journey regarding motivation and trust in me and this important topic. Therefore,
I am grateful to DBU for their fnancial support, very positive encouragement and for providing
access to the DBU network.
A special thanks goes to my colleagues at ASG Analytik-Service AG for our discussions, the
support, freedom and trust they gave me regarding many experimental investigations and pre-
tests, and their very fast and competent support for problem solving. Without them, this journey
would not have come this far.
I would like to thank ChemCom Industries B.V. for the fnancial support and motivation to
realize the investigations on this new process concept.
Last but not least, I would like to thank my partner Mafalda as well as my family and friends
for their patience, support and care during the past years. Thank you for being who you are and
helping me to get where I am today.
117
(a) (b)
(c) (d)
(e) (f)
(g)
Figure A.1:
OME synthesis from MeOH-pFA over (a) A15, (b) A36, (c) A46, (d) Dowex, (e)
H-BEA25, (f) H-MFI90, (g) Nafon (conditions: pFA/MeOH = 1
.
50
gg
�1
, cata-
lyst/(MeOH+pFA) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch). The termination time represents
the time after which 90 % of the OME5 concentration after 24 h are obtained.
Table A.1:
OME synthesis from MeOH-pFA over A15 (conditions: pFA/MeOH = 1
.
50
g g�1
, A15/(MeOH+pFA) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch).
The concentrations are presented in mass fractions and the duration in minutes.
t in min 0 0 1 5 10 15 20 30 45 60 75 90 120 150 180 240 300 1440
FA 0.5862 0.5653 0.5436 0.5143 0.465 0.4242 0.4182 0.3924 0.3684 0.3681 0.359 0.3537 0.3602 0.3469 0.3402 0.3446 0.3512 0.3431
H2O 0.0238 0.0234 0.0253 0.0315 0.0419 0.0534 0.0557 0.0671 0.0784 0.0805 0.0976 0.0927 0.0944 0.0952 0.0879 0.0868 0.091 0.0974
MeOH 0.3871 0.4076 0.4261 0.4319 0.4218 0.4129 0.3747 0.3479 0.3434 0.269 0.2455 0.2391 0.2202 0.2264 0.2156 0.2113 0.208 0.2053
OME1 0.0009 0.0013 0.0013 0.0091 0.0301 0.0461 0.0633 0.0799 0.0656 0.1144 0.1189 0.1243 0.1264 0.1122 0.1338 0.133 0.1277 0.1208
OME2 0.0007 0.001 0.0013 0.0066 0.0216 0.033 0.0454 0.0574 0.0704 0.082 0.0855 0.089 0.0901 0.0934 0.0956 0.0938 0.0896 0.0903
OME3 0.0004 0.0005 0.0006 0.0029 0.0098 0.0151 0.0212 0.0271 0.0357 0.041 0.0439 0.0469 0.0489 0.0555 0.0547 0.0547 0.0534 0.0547
OME4 0.0003 0.0003 0.0003 0.0015 0.0051 0.0079 0.0113 0.0146 0.0197 0.0229 0.0251 0.0272 0.0293 0.0342 0.0343 0.0352 0.0354 0.0367
OME5 0.0002 0.0001 0.0002 0.0007 0.0023 0.0035 0.005 0.0066 0.009 0.0106 0.0118 0.0129 0.0143 0.0169 0.0173 0.0182 0.0189 0.02
OME6 0.0001 0.0001 0.0001 0.0003 0.0011 0.0017 0.0025 0.0033 0.0045 0.0054 0.006 0.0067 0.0075 0.009 0.0093 0.01 0.0107 0.0118
OME7 0.0001 0 0.0006 0.0002 0.0006 0.0009 0.0012 0.0016 0.0022 0.0027 0.003 0.0034 0.0039 0.0046 0.0049 0.0053 0.0058 0.0067
OME8 0 0 0.0003 0.0005 0.0003 0.0004 0.0006 0.0008 0.0011 0.0014 0.0015 0.0017 0.002 0.0024 0.0025 0.0028 0.0031 0.0038
OME9 0 0 0 0 0 0.0002 0.0003 0.0004 0.0006 0.0007 0.0008 0.0009 0.001 0.0012 0.0013 0.0015 0.0017 0.0022
OME10 0 0 0 0 0 0 0 0 0.0003 0.0004 0.0004 0.0005 0.0005 0.0007 0.0007 0.0008 0.0009 0.0013
OME11 000000000000000000.0007
TRI 0 0 0 0 0 0 0 0.0003 0.0004 0.0005 0.0006 0.0008 0.001 0.0012 0.0014 0.0018 0.0023 0.0049
MEFO 0.0004 0.0004 0.0004 0.0005 0.0006 0.0006 0.0006 0.0006 0.0003 0.0005 0.0005 0.0004 0.0004 0.0002 0.0004 0.0004 0.0003 0.0003
Table A.2:
OME synthesis from MeOH-pFA over A36 (conditions: pFA/MeOH = 1
.
53
g g�1
, A36/(MeOH+pFA) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch).
The concentrations are presented in mass fractions and the duration in minutes.
t in min 0 0 1 5 10 15 20 30 45 60 75 90 120 150 180 240 300 1440
FA 0.5837 0.5892 0.5457 0.5197 0.4905 0.4692 0.4521 0.4296 0.4057 0.3921 0.385 0.3783 0.3714 0.3631 0.3623 0.3571 0.354 0.3452
H2O 0.0236 0.0255 0.0356 0.0468 0.0553 0.0646 0.0665 0.0697 0.0776 0.078 0.0826 0.0828 0.0853 0.0849 0.0858 0.0861 0.0901 0.0861
MeOH 0.3902 0.383 0.3733 0.3563 0.3375 0.3096 0.293 0.2657 0.2378 0.2197 0.2012 0.201 0.1942 0.1916 0.1868 0.188 0.187 0.1639
OME1 0.0007 0.0006 0.0209 0.0355 0.0521 0.0696 0.0829 0.1013 0.1181 0.1281 0.1342 0.1353 0.1361 0.1379 0.1376 0.1375 0.1324 0.145
OME2 0.0006 0.0005 0.0138 0.0235 0.0361 0.0476 0.057 0.0703 0.0823 0.0904 0.0951 0.0956 0.0966 0.0978 0.0975 0.0963 0.0957 0.1016
OME3 0.0003 0.0002 0.0058 0.0099 0.0157 0.0213 0.026 0.0331 0.0402 0.0459 0.0498 0.0514 0.0541 0.0563 0.0573 0.0573 0.0582 0.0604
OME4 0.0001 0.0001 0.0025 0.0046 0.0074 0.0103 0.0129 0.017 0.0213 0.025 0.0279 0.0294 0.0322 0.0346 0.0361 0.0371 0.0385 0.0398
OME5 0.0001 0.0001 0.0009 0.0017 0.0028 0.004 0.0051 0.0069 0.0089 0.0108 0.0123 0.0132 0.0149 0.0165 0.0175 0.0187 0.0197 0.0211
OME6 0 0 0.0004 0.0007 0.0012 0.0017 0.0022 0.003 0.004 0.0049 0.0057 0.0062 0.0072 0.0081 0.0088 0.0098 0.0106 0.0119
OME7 0 0 0.0002 0.0003 0.0005 0.0007 0.001 0.0013 0.0018 0.0022 0.0026 0.0028 0.0034 0.0039 0.0043 0.0049 0.0054 0.0066
OME8 0 0 0 0.0002 0.0002 0.0003 0.0004 0.0006 0.0008 0.001 0.0012 0.0013 0.0016 0.0019 0.0021 0.0024 0.0027 0.0036
OME9 0 0 0 0 0 0.0002 0.0002 0.0003 0.0004 0.0005 0.0006 0.0007 0.0008 0.0009 0.0011 0.0013 0.0014 0.0021
OME10 0 0 0 0 0 0 0 0 0 0 0.0003 0.0003 0.0004 0.0005 0.0006 0.0007 0.0008 0.0012
OME11 000000000000000000.0008
TRI 0 0 0 0 0 0 0 0.0004 0.0005 0.0007 0.0008 0.001 0.0012 0.0015 0.0018 0.0023 0.0029 0.0102
MEFO 0.0008 0.0007 0.0008 0.0008 0.0008 0.0008 0.0008 0.0008 0.0008 0.0008 0.0008 0.0007 0.0007 0.0006 0.0006 0.0006 0.0006 0.0007
Table A.3:
OME synthesis from MeOH-pFA over A46 (conditions: pFA/MeOH = 1
.
50
g g�1
, A46/(MeOH+pFA) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch).
The concentrations are presented in mass fractions and the duration in minutes.
t in min 0 0 1 5 10 15 20 30 45 60 75 90 120 150 180 240 300 1440
FA 0.5711 0.5675 0.5508 0.5157 0.4828 0.4562 0.4237 0.411 0.3857 0.376 0.3617 0.3549 0.3444 0.3453 0.3426 0.3399 0.3373 0.325
H2O 0.0219 0.0265 0.0294 0.0373 0.0492 0.062 0.0666 0.0748 0.0857 0.087 0.0875 0.0842 0.0825 0.0701 0.0961 0.0938 0.0955 0.0869
MeOH 0.4046 0.403 0.4053 0.3925 0.3725 0.3466 0.3332 0.2879 0.2664 0.2279 0.2327 0.2247 0.158 0.1512 0.1553 0.1536 0.1512 0.148
OME1 0.0007 0.0008 0.0059 0.0227 0.0397 0.055 0.0722 0.0917 0.1057 0.121 0.125 0.1305 0.1572 0.1626 0.1555 0.1563 0.1563 0.1623
OME2 0.0007 0.0008 0.0046 0.0172 0.0299 0.042 0.0543 0.0686 0.0779 0.0902 0.0915 0.0954 0.1086 0.1168 0.1085 0.1082 0.1078 0.1113
OME3 0.0003 0.0004 0.002 0.0077 0.0137 0.0197 0.0259 0.0337 0.0393 0.0476 0.0488 0.052 0.0663 0.0677 0.0628 0.0635 0.0636 0.0661
OME4 0.0001 0.0003 0.0009 0.0037 0.0068 0.0101 0.0134 0.0179 0.0215 0.0269 0.028 0.0305 0.0412 0.0427 0.0395 0.0411 0.0418 0.044
OME5 0.0001 0.0002 0.0004 0.0015 0.0028 0.0042 0.0056 0.0076 0.0094 0.0121 0.0128 0.0142 0.0204 0.0214 0.0195 0.0209 0.0218 0.0237
OME6 0.0001 0.0001 0.0002 0.0007 0.0013 0.0019 0.0025 0.0035 0.0044 0.0058 0.0062 0.007 0.0108 0.0112 0.0101 0.0112 0.012 0.0137
OME7 0 0.0001 0.0001 0.0003 0.0006 0.0009 0.0011 0.0016 0.002 0.0027 0.0029 0.0033 0.0054 0.0055 0.005 0.0057 0.0062 0.0076
OME8 0 0 0 0.0001 0.0003 0.0004 0.0005 0.0007 0.001 0.0013 0.0014 0.0016 0.0026 0.0027 0.0025 0.0029 0.0032 0.0042
OME9 0 0 0 0 0 0.0002 0.0003 0.0004 0.0005 0.0006 0.0007 0.0008 0.0013 0.0014 0.0013 0.0015 0.0017 0.0023
OME10 0 0 0 0 0 0 0 0 0 0.0003 0.0004 0.0004 0.0006 0.0007 0.0007 0.0008 0.0009 0.0013
OME11 000000000000000000.0008
TRI 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0.002
MEFO 0.0005 0.0005 0.0005 0.0006 0.0006 0.0006 0.0007 0.0007 0.0006 0.0006 0.0006 0.0006 0.0007 0.0007 0.0006 0.0006 0.0006 0.0007
Table A.4:
OME synthesis from MeOH-pFA over Dowex (conditions: pFA/MeOH = 1
.
50
g g�1
, Dowex/(MeOH+pFA) = 1
.
0
wt
%, 60
◦C
,8
bar
,
batch). The concentrations are presented in mass fractions and the duration in minutes.
t in min 0 0 1 5 10 15 20 30 45 60 75 90 120 150 180 240 300
FA 0.5652 0.5591 0.4365 0.3962 0.3728 0.3622 0.3564 0.3502 0.3466 0.3431 0.3426 0.3399 0.3396 0.3372 0.3376 0.3382 0.337
H2O 0.0255 0.0257 0.069 0.0799 0.0811 0.0854 0.0854 0.0887 0.087 0.0916 0.0916 0.0901 0.0909 0.0908 0.0887 0.0889 0.0856
MeOH 0.3906 0.3814 0.1687 0.1441 0.1227 0.1033 0.1082 0.1065 0.1077 0.1011 0.1018 0.1059 0.1146 0.1049 0.1055 0.1066 0.1068
OME1 0.0056 0.0108 0.1203 0.1453 0.1594 0.1684 0.1674 0.1687 0.169 0.1695 0.1685 0.1691 0.1665 0.1723 0.1706 0.1692 0.1706
OME2 0.0054 0.0098 0.0994 0.1138 0.1225 0.1269 0.1251 0.1232 0.1214 0.1205 0.1191 0.1184 0.1162 0.1191 0.1189 0.1183 0.1194
OME3 0.0029 0.0052 0.0512 0.0581 0.0647 0.0689 0.0694 0.07 0.0706 0.0713 0.0711 0.0707 0.0687 0.0704 0.0708 0.0706 0.0713
OME4 0.0015 0.0029 0.0281 0.0325 0.0375 0.0409 0.042 0.0435 0.045 0.0465 0.047 0.047 0.0454 0.0464 0.0472 0.0471 0.0475
OME5 0.0007 0.0014 0.0132 0.0151 0.0177 0.0197 0.0205 0.0217 0.023 0.0243 0.025 0.0251 0.0243 0.0246 0.0253 0.0253 0.0254
OME6 0.0004 0.0008 0.0067 0.0077 0.0091 0.0103 0.0108 0.0117 0.0126 0.0137 0.0143 0.0144 0.0142 0.0141 0.0146 0.0147 0.0148
OME7 0.0002 0.0004 0.0032 0.0036 0.0043 0.0048 0.0051 0.0056 0.0061 0.0068 0.0073 0.0074 0.0077 0.0073 0.0076 0.0077 0.0077
OME8 0 0.0002 0.0015 0.0016 0.0019 0.0022 0.0023 0.0025 0.0028 0.0031 0.0034 0.0035 0.0039 0.0036 0.0037 0.0038 0.0038
OME9 0 0 0 0 0 0.001 0.001 0.0011 0.0013 0.0015 0.0016 0.0017 0.0021 0.0018 0.0019 0.0019 0.0019
OME10 0 0 0 0 0 0 0 0 0.0005 0.0007 0.0007 0.0008 0.001 0.0009 0.0009 0.0009 0.0009
OME11 00000000000000000
TRI 0 0 0 0 0 0 0 0 0 0 0 0 0 0.0005 0.0006 0.0009 0.0012
MEFO 0.0022 0.0021 0.0023 0.0023 0.0062 0.006 0.0063 0.0065 0.0064 0.0063 0.006 0.0061 0.005 0.0061 0.0062 0.0061 0.0062
Table A.5:
OME synthesis from MeOH-pFA over H-BEA 25 (conditions: pFA/MeOH = 1
.
50
g g�1
, H-BEA 25/(MeOH+pFA) = 1
.
0
wt
%, 60
◦C
,
8 bar, batch). The concentrations are presented in mass fractions and the duration in minutes.
t in min 0 0 1 5 10 15 20 30 45 60 75 90 120 150 180 240 300 1440
FA 0.5911 0.5873 0.5959 0.5846 0.5755 0.5742 0.5619 0.5604 0.5487 0.5323 0.5258 0.5109 0.4937 0.4846 0.4681 0.4423 0.4228 0.3509
H2O 0.0218 0.0243 0.0242 0.0281 0.0273 0.0324 0.0331 0.0366 0.0411 0.0426 0.0426 0.0491 0.0491 0.0513 0.0592 0.0673 0.0737 0.091
MeOH 0.3839 0.3814 0.3744 0.3764 0.3793 0.3682 0.3631 0.3567 0.3486 0.3462 0.3389 0.3294 0.3167 0.3041 0.2909 0.2484 0.2444 0.154
OME1 0.001 0.0023 0.0019 0.0039 0.0064 0.009 0.0148 0.0165 0.0219 0.0278 0.0345 0.0408 0.0477 0.0521 0.0626 0.0869 0.0951 0.1552
OME2 0.0007 0.0017 0.0015 0.0033 0.0057 0.008 0.0134 0.0149 0.0197 0.0252 0.0302 0.0355 0.043 0.0473 0.0534 0.071 0.0744 0.105
OME3 0.0003 0.0008 0.0006 0.0013 0.0023 0.0033 0.0056 0.0062 0.0084 0.0109 0.0125 0.0151 0.0205 0.0237 0.0262 0.0345 0.0366 0.0585
OME4 0.0001 0.0004 0.0003 0.0007 0.0013 0.0019 0.0033 0.0037 0.0051 0.0067 0.0072 0.0089 0.0133 0.0162 0.0175 0.0224 0.0237 0.0376
OME5 0.0001 0.0002 0.0001 0.0003 0.0006 0.001 0.0017 0.0019 0.0026 0.0034 0.0036 0.0045 0.0071 0.009 0.0097 0.012 0.0127 0.0199
OME6 0.0001 0.0003 0.0001 0.0002 0.0003 0.0005 0.0009 0.001 0.0014 0.0018 0.0018 0.0023 0.0039 0.0051 0.0055 0.0066 0.007 0.0114
OME7 0 0.0001 0 0.0001 0.0002 0.0002 0.0004 0.0005 0.0006 0.0009 0.0009 0.0011 0.0019 0.0026 0.0028 0.0034 0.0038 0.0062
OME8 0 0 0 0 0.0001 0.0001 0.0002 0.0002 0.0003 0.0004 0.0005 0.0006 0.001 0.0014 0.0015 0.0018 0.0021 0.0034
OME9 0 0 0 0 0 0 0.0001 0.0001 0.0002 0.0002 0.0002 0.0003 0.0005 0.0008 0.0009 0.001 0.0012 0.0019
OME10 0 0 0 0 0 0 0 0 0 0 0 0.0002 0.0003 0.0005 0.0005 0.0006 0.0007 0.0011
OME11 000000000000000000.0007
TRI 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0.0003 0.0013
MEFO 0.001 0.0011 0.0011 0.0011 0.0012 0.0011 0.0015 0.0013 0.0015 0.0016 0.0015 0.0014 0.0014 0.0013 0.0014 0.0017 0.0016 0.002
Table A.6:
OME synthesis from MeOH-pFA over H-MFI 90 (conditions: pFA/MeOH = 1
.
50
g g�1
, H-MFI 90/(MeOH+pFA) = 1
.
0
wt
%, 60
◦C
,
8 bar, batch). The concentrations are presented in mass fractions and the duration in minutes.
t in min 0 0 1 5 10 15 20 30 45 60 75 90 120 150 180 240 300 1440
FA 0.5592 0.5558 0.5543 0.5563 0.5468 0.5386 0.5347 0.5222 0.504 0.4697 0.4688 0.4632 0.4347 0.4181 0.4048 0.3893 0.3608 0.2755
H2O 0.0237 0.0269 0.0292 0.0286 0.0318 0.0314 0.0363 0.0362 0.0434 0.046 0.0527 0.0529 0.0603 0.0641 0.0719 0.0764 0.0836 0.0915
MeOH 0.4151 0.4144 0.4121 0.4068 0.4069 0.4052 0.3975 0.3955 0.3831 0.3876 0.3599 0.3448 0.3254 0.2971 0.2732 0.2309 0.1832 0.1343
OME1 0.0006 0.0007 0.0014 0.003 0.0055 0.0095 0.0121 0.0178 0.0268 0.0372 0.0459 0.0541 0.0699 0.0837 0.0949 0.1141 0.1409 0.2001
OME2 0.0004 0.0006 0.0011 0.0023 0.0042 0.0074 0.0094 0.0138 0.0207 0.0287 0.0351 0.041 0.0525 0.0625 0.0698 0.0824 0.1002 0.1269
OME3 0.0002 0.0004 0.0004 0.0009 0.0016 0.0028 0.0036 0.0054 0.0082 0.0116 0.0144 0.017 0.0221 0.0269 0.0307 0.0375 0.0475 0.0694
OME4 0.0001 0.0002 0.0003 0.0006 0.0012 0.0021 0.0027 0.004 0.0062 0.0088 0.0109 0.0127 0.0164 0.0199 0.0224 0.0269 0.0343 0.0422
OME5 0 0.0001 0.0001 0.0003 0.0006 0.001 0.0013 0.0019 0.003 0.0043 0.0053 0.0062 0.008 0.0097 0.0111 0.0134 0.0177 0.0206
OME6 0 0.0001 0.0001 0.0002 0.0003 0.0005 0.0006 0.0009 0.0014 0.0021 0.0026 0.003 0.0039 0.0048 0.0056 0.0068 0.0096 0.0108
OME7 0 0 0 0.0001 0.0002 0.0003 0.0004 0.0005 0.0008 0.0012 0.0014 0.0017 0.0022 0.0027 0.0031 0.0037 0.0053 0.0054
OME8 0 0 0 0 0 0.0001 0.0002 0.0003 0.0004 0.0006 0.0007 0.0009 0.0011 0.0014 0.0016 0.002 0.0028 0.0028
OME9 0 0 0 0 0 0 0 0.0001 0.0002 0.0003 0.0004 0.0005 0.0006 0.0007 0.0009 0.0011 0.0015 0.0015
OME10 0 0 0 0 0 0 0 0 0 0 0 0 0.0003 0.0004 0.0005 0.0006 0.0009 0.0008
OME11 000000000000000000
TRI 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0.0017
MEFO 0.0008 0.0008 0.0009 0.0009 0.001 0.0011 0.0012 0.0013 0.0016 0.0018 0.002 0.0022 0.0026 0.0079 0.0096 0.0149 0.0119 0.0165
Table A.7:
OME synthesis from MeOH-pFA over Nafon (conditions: pFA/MeOH = 1
.
49
g g�1
, Nafon/(MeOH+pFA) = 1
.
0
wt
%, 60
◦C
,8
bar
,
batch). The concentrations are presented in mass fractions and the duration in minutes.
t in min 0 0 1 5 10 15 20 30 45 60 75 90 120 150 180 240 300 1440
FA 0.583 0.5778 0.5786 0.5709 0.5664 0.554 0.5426 0.5309 0.513 0.4961 0.4802 0.4664 0.4455 0.4224 0.4091 0.3936 0.3819 0.353
H2O 0.0233 0.0262 0.0262 0.029 0.0296 0.0351 0.0372 0.0434 0.0454 0.0543 0.0571 0.0599 0.0639 0.0722 0.0756 0.0792 0.0823 0.0841
MeOH 0.3917 0.3881 0.3866 0.3844 0.3769 0.3739 0.3679 0.3548 0.3455 0.3275 0.3084 0.2932 0.2697 0.2447 0.236 0.2065 0.1942 0.1716
OME1 0.0005 0.0031 0.0034 0.0067 0.0118 0.0159 0.0224 0.0301 0.0404 0.0509 0.0637 0.0738 0.089 0.104 0.1101 0.1246 0.1307 0.1413
OME2 0.0005 0.0025 0.0027 0.005 0.0085 0.0116 0.0163 0.0221 0.0296 0.0373 0.0468 0.0544 0.0658 0.0769 0.0817 0.0918 0.0961 0.1006
OME3 0.0002 0.001 0.0011 0.002 0.0036 0.0051 0.0073 0.0101 0.0139 0.0179 0.0228 0.0269 0.0335 0.0398 0.0431 0.0497 0.0534 0.0603
OME4 0.0001 0.0005 0.0004 0.0009 0.0016 0.0023 0.0033 0.0048 0.0068 0.0089 0.0116 0.014 0.0179 0.0218 0.024 0.0286 0.0315 0.0398
OME5 0 0.0002 0.0002 0.0003 0.0006 0.0009 0.0013 0.0018 0.0027 0.0036 0.0048 0.0058 0.0076 0.0095 0.0106 0.013 0.0147 0.021
OME6 0 0.0001 0.0001 0.0001 0.0003 0.0004 0.0006 0.0008 0.0012 0.0016 0.0021 0.0026 0.0035 0.0044 0.005 0.0063 0.0072 0.0118
OME7 0 0 0 0 0.0001 0.0002 0.0002 0.0004 0.0005 0.0007 0.001 0.0012 0.0016 0.002 0.0024 0.003 0.0035 0.0064
OME8 0 0 0 0 0 0 0.0001 0.0002 0.0002 0.0003 0.0005 0.0006 0.0008 0.001 0.0011 0.0015 0.0017 0.0035
OME9 0 0 0 0 0 0 0 0 0 0.0002 0.0002 0.0003 0.0004 0.0005 0.0006 0.0008 0.0009 0.002
OME10 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0.0004 0.0005 0.0012
OME11 000000000000000000.0007
TRI 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0.0004 0.0005 0.0022
MEFO 0.0007 0.0006 0.0006 0.0007 0.0007 0.0007 0.0008 0.0007 0.0008 0.0008 0.0008 0.0008 0.0008 0.0008 0.0008 0.0008 0.0007 0.0006
(a) (b)
(c) (d)
(e) (f)
(g)
Figure A.2:
OME synthesis from OME
1
-TRI over (a) A15, (b) A36, (c) A46, (d) Dowex,
(e) H-BEA25, (f) H-MFI90, (g) Nafon (conditions: OME
1
/TRI = 2
.
00
gg
�1
,
catalyst/(OME
1
+TRI) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch). The termination time
represents the time after which 90 % of the OME
5
concentration after 24
h
are
obtained.
Table A.8:
OME synthesis from OME
1
-TRI over A15 (conditions: OME
1
/TRI = 1
.
99
g g�1
, A15/(OME
1
+TRI) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch).
The concentrations are presented in mass fractions and the duration in minutes.
t in min 0 0 1 5 10 15 20 30 45 60 75 90 120 150 180 240 300 1440
FA 0.0059 0 0.0024 0.0136 0.0154 0.0218 0.0186 0.02 0.0194 0.0182 0.0203 0.0203 0.0167 0.0168 0.0172 0.0168 0.0166 0.0164
H2O 000000000000000000
MeOH 0.0026 0.0005 0.0062 0.0089 0.0102 0.0105 0.0103 0.01 0.0094 0.0096 0.0086 0.0084 0.0083 0.008 0.0078 0.0076 0.0072 0.0064
OME1 0.5836 0.5878 0.5683 0.4521 0.3535 0.3179 0.3075 0.3009 0.2967 0.2684 0.2919 0.2905 0.2905 0.2897 0.288 0.2891 0.2888 0.2847
OME2 0.002 0.0035 0.0197 0.1419 0.2092 0.217 0.2187 0.2191 0.2172 0.2224 0.2157 0.2159 0.2164 0.2164 0.2164 0.2162 0.216 0.2141
OME3 0.001 0.0006 0.0019 0.0518 0.1126 0.1322 0.1389 0.1418 0.1416 0.1473 0.1417 0.1421 0.1429 0.1432 0.1433 0.1432 0.1432 0.1424
OME4 0.001 0.0009 0.0008 0.0241 0.0697 0.0908 0.0994 0.1034 0.104 0.1089 0.1047 0.1053 0.1062 0.1065 0.1066 0.1066 0.1067 0.1064
OME5 0.0005 0.0001 0.0001 0.0095 0.0357 0.0513 0.0586 0.0623 0.0631 0.0666 0.064 0.0645 0.0653 0.0656 0.0656 0.0657 0.0658 0.0658
OME6 0.0003 0 0 0.0044 0.0202 0.0318 0.0379 0.0412 0.0421 0.0447 0.043 0.0434 0.0441 0.0443 0.0444 0.0445 0.0445 0.0446
OME7 0.0002 0 0 0.0021 0.0114 0.0193 0.024 0.0267 0.0276 0.0294 0.0283 0.0286 0.0291 0.0293 0.0294 0.0295 0.0295 0.0295
OME8 0.0001 0 0 0.001 0.0065 0.0118 0.0152 0.0174 0.0181 0.0195 0.0187 0.019 0.0194 0.0195 0.0196 0.0196 0.0197 0.0196
OME9 0 0 0 0 0 0.0073 0.0098 0.0115 0.0121 0.0131 0.0126 0.0128 0.013 0.0132 0.0132 0.0133 0.0133 0.0133
OME10 0 0 0 0 0 0 0 0 0.0083 0.009 0.0087 0.0088 0.009 0.0091 0.0091 0.0092 0.0092 0.0092
OME11 0 0 0 0 0 0 0 0 0 0.0065 0.0062 0.0063 0.0065 0.0066 0.0066 0.0066 0.0067 0.0067
TRI 0.4028 0.4067 0.4005 0.2908 0.1557 0.0883 0.0611 0.0457 0.0403 0.0363 0.0354 0.0336 0.032 0.0312 0.0318 0.0309 0.031 0.0316
MEFO 0 0 0 0 0 0 0 0 0.0001 0.0002 0.0003 0.0003 0.0006 0.0008 0.0009 0.0012 0.0017 0.0092
Table A.9:
OME synthesis from OME
1
-TRI over A36 (conditions: OME
1
/TRI = 2
.
00
g g�1
, A36/(OME
1
+TRI) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch).
The concentrations are presented in mass fractions and the duration in minutes.
t in min 0 0 1 5 10 15 20 30 45 60 75 90 120 150 180 240 300 1440
FA 0 0 0 0.0076 0.0116 0.0183 0.0221 0.0245 0.0265 0.0263 0.0256 0.0252 0.0241 0.0237 0.0238 0.0224 0.0222 0.0217
H2O 0 0 0.0018 0.0013 0.0009 0.0007 0.0006 0 0 0 0 0 0 0 0 0 0 0
MeOH 0 0.0058 0.0003 0.0095 0.0107 0.0112 0.0104 0.0106 0.01 0.0094 0.0089 0.0098 0.0091 0.0087 0.0084 0.0079 0.0076 0.0058
OME1 0.6621 0.6456 0.654 0.6064 0.5362 0.461 0.4144 0.3637 0.338 0.3296 0.3255 0.3209 0.3191 0.3175 0.3169 0.3134 0.3141 0.31
OME2 0 0.0061 0.0011 0.0503 0.1232 0.1815 0.2066 0.2226 0.2254 0.2251 0.2259 0.223 0.2243 0.2254 0.226 0.2274 0.2275 0.2256
OME3 0 0.0005 0.0004 0.0092 0.0363 0.0723 0.0976 0.1245 0.136 0.1394 0.1414 0.1404 0.1422 0.1432 0.1438 0.1461 0.1462 0.1454
OME4 0 0.0002 0.0006 0.0025 0.0133 0.0336 0.0523 0.0775 0.0912 0.096 0.0984 0.0986 0.1003 0.1013 0.1018 0.1041 0.1043 0.1041
OME5 0 0 0.0003 0.0006 0.0041 0.0128 0.0226 0.0389 0.0495 0.0537 0.0557 0.0568 0.0579 0.0586 0.0588 0.0606 0.0608 0.0609
OME6 0 0 0.0002 0.0002 0.0015 0.0053 0.0105 0.0209 0.0289 0.0323 0.0339 0.0356 0.0362 0.0367 0.0368 0.0381 0.0383 0.0387
OME7 0 0 0.0001 0 0.0006 0.0023 0.0049 0.0109 0.0162 0.0187 0.0199 0.0215 0.0219 0.0222 0.0222 0.0231 0.0233 0.0239
OME8 0 0 0 0 0.0003 0.001 0.0023 0.0057 0.0091 0.0108 0.0116 0.0129 0.0131 0.0133 0.0133 0.0139 0.0141 0.0145
OME9 0 0 0 0 0 0.0005 0.0012 0.0031 0.0053 0.0065 0.0071 0.008 0.0082 0.0084 0.0084 0.0088 0.0089 0.0092
OME10 0 0 0 0 0 0 0.0006 0.0017 0.0031 0.0039 0.0043 0.0049 0.005 0.0051 0.0052 0.0054 0.0055 0.0057
OME11 0 0 0 0 0 0 0 0.0011 0.002 0.0026 0.0029 0.0031 0.0034 0.0034 0.0035 0.0037 0.0037 0.0038
TRI 0.3379 0.3416 0.3412 0.3124 0.2614 0.1994 0.1538 0.0943 0.0588 0.0455 0.0387 0.0389 0.0348 0.0318 0.0307 0.0242 0.0224 0.0235
MEFO 0 0 0 0 0 0 0 0 0 0.0002 0.0002 0.0003 0.0004 0.0005 0.0006 0.0008 0.0011 0.0073
Table A.10:
OME synthesis from OME
1
-TRI over A46 (conditions: OME
1
/TRI = 2
.
00
g g�1
, A46/(OME
1
+TRI) = 1
.
0
wt
%, 60
◦C
,8
bar
, batch).
The concentrations are presented in mass fractions and the duration in minutes.
t in min 0 0 1 5 10 15 20 30 45 60 75 90 120 150 180 240 300 1440
FA 0.0024 0.0025 0.0028 0.0046 0.0056 0.0065 0.0066 0.0083 0.0049 0.005 0.0045 0.005 0.0035 0.0059 0.0045 0.0045 0.0045 0.0044
H2O 000000000000000000
MeOH 0 0.0003 0.0041 0.0048 0.005 0.0049 0.0046 0.0044 0.0041 0.004 0.004 0.0038 0.0037 0.0036 0.0036 0.0035 0.0035 0.0034
OME1 0.6961 0.6974 0.6792 0.4957 0.3945 0.3571 0.3528 0.3421 0.3364 0.3305 0.3273 0.3261 0.3263 0.3253 0.3226 0.3233 0.3222 0.3223
OME2 0.0001 0.0004 0.0134 0.1811 0.2297 0.234 0.2308 0.2321 0.2326 0.232 0.2316 0.232 0.2318 0.2312 0.2326 0.2321 0.2321 0.233
OME3 0 0.0001 0.0009 0.0722 0.1231 0.1395 0.1408 0.1445 0.1464 0.1476 0.148 0.1482 0.1486 0.1483 0.1496 0.1493 0.1496 0.1503
OME4 0 0.0002 0.0003 0.0352 0.0747 0.0928 0.0959 0.1003 0.1026 0.1043 0.1051 0.1053 0.1058 0.1057 0.1068 0.1067 0.1071 0.1076
OME5 0.0001 0 0 0.0134 0.0361 0.0495 0.0529 0.0565 0.0584 0.06 0.0608 0.0609 0.0614 0.0613 0.0622 0.0621 0.0624 0.0627
OME6 0 0 0 0.0057 0.0187 0.0284 0.0313 0.0342 0.0358 0.0371 0.0379 0.0379 0.0383 0.0383 0.0389 0.039 0.0391 0.0392
OME7 0 0 0 0.0024 0.0093 0.0155 0.0177 0.0199 0.021 0.022 0.0226 0.0226 0.0229 0.0229 0.0234 0.0234 0.0236 0.0235
OME8 0 0 0 0.0011 0.0047 0.0085 0.01 0.0115 0.0123 0.0129 0.0134 0.0134 0.0136 0.0136 0.014 0.014 0.0141 0.0139
OME9 0 0 0 0 0 0.0048 0.0058 0.0069 0.0074 0.0079 0.0083 0.0083 0.0084 0.0084 0.0087 0.0087 0.0088 0.0086
OME10 0 0 0 0 0 0 0 0 0.0046 0.0049 0.0051 0.0051 0.0052 0.0053 0.0054 0.0054 0.0055 0.0054
OME11 0 0 0 0 0 0 0 0 0 0.0033 0.0035 0.0035 0.0036 0.0036 0.0037 0.0037 0.0037 0.0037
TRI 0.3012 0.2992 0.2992 0.1839 0.0986 0.0585 0.0507 0.0395 0.0336 0.0285 0.028 0.0279 0.0267 0.0264 0.0238 0.0239 0.0234 0.0209
MEFO 0 0 0 0 0 0 0 0 0 0 0 0 0.0002 0.0002 0.0002 0.0003 0.0004 0.0012
Table A.11:
OME synthesis from OME
1
-TRI over Dowex (conditions: OME
1
/TRI = 2
.
01
g g�1
, Dowex/(OME
1
+TRI) = 1
.
0
wt
%, 60
◦C
,8
bar
,
batch). The concentrations are presented in mass fractions and the duration in minutes.
t in min 0 0 1 5 10 15 20 30 45 60 90 120 180 240 300 1440
FA 0.0024 0.0024 0.003 0.0055 0.016 0.0233 0.0265 0.0229 0.0254 0.0154 0.0102 0.0158 0.019 0.019 0.019 0.0181
H2O 0000000000000000
MeOH 0 0.0009 0.0062 0.0067 0.0067 0.0066 0.0062 0.0057 0.0053 0.0051 0.0049 0.0048 0.0043 0.0042 0.0042 0.0042
OME1 0.6245 0.647 0.6212 0.5667 0.4229 0.3499 0.3358 0.3225 0.3067 0.3071 0.3034 0.291 0.2984 0.2948 0.2959 0.2836
OME2 0 0.0017 0.0087 0.0805 0.2066 0.23 0.2287 0.2295 0.2287 0.2294 0.2309 0.2289 0.2284 0.2284 0.2283 0.2276
OME3 0 0.0003 0.0007 0.016 0.0884 0.13 0.1366 0.1439 0.1485 0.1503 0.1529 0.1548 0.1524 0.1534 0.1531 0.1558
OME4 0 0.0001 0.0002 0.0043 0.041 0.0793 0.0884 0.0975 0.1043 0.1068 0.1098 0.1123 0.1102 0.1113 0.111 0.1138
OME5 0 0.0001 0 0.0011 0.015 0.0386 0.0461 0.0536 0.0593 0.0617 0.064 0.0658 0.0647 0.0654 0.0652 0.0672
OME6 0 0 0 0.0005 0.006 0.0201 0.0257 0.0315 0.0363 0.0383 0.0402 0.0415 0.0409 0.0414 0.0413 0.0428
OME7 0 0 0 0.0002 0.0024 0.01 0.0137 0.0178 0.0213 0.0229 0.0243 0.0253 0.0249 0.0253 0.0253 0.0263
OME8 0 0 0 0 0.001 0.0049 0.0072 0.01 0.0124 0.0136 0.0146 0.0152 0.0151 0.0153 0.0153 0.0159
OME9 0 0 0 0 0 0.0025 0.0039 0.0057 0.0073 0.0081 0.0089 0.0093 0.0092 0.0094 0.0094 0.0098
OME10 0 0 0 0 0 0 0 0 0.0043 0.0048 0.0053 0.0056 0.0056 0.0057 0.0057 0.0059
OME11 0 0 0 0 0 0 0 0 0 0.003 0.0034 0.0036 0.0036 0.0037 0.0036 0.0037
TRI 0.3731 0.3474 0.3595 0.3178 0.1941 0.1047 0.0812 0.0593 0.0402 0.0335 0.0274 0.0261 0.0229 0.0223 0.0224 0.0233
MEFO 0 0 0.0004 0.0006 0 0 0 0 0 0 0 0 0.0002 0.0003 0.0004 0.0019
Table A.12:
OME synthesis from OME
1
-TRI over H-BEA 25 (conditions: OME
1
/TRI = 2
.
01
g g�1
, H-BEA 25/(OME
1
+TRI) = 1
.
0
wt
%, 60
◦C
,
8 bar, batch). The concentrations are presented in mass fractions and the duration in minutes.
t in min 0 0 1 5 10 15 20 30 45 60 75 90 120 150 180 240 300 1440
FA 0.0018 0.0028 0.0023 0.0023 0.0036 0.0044 0.0057 0.0054 0.0064 0.0076 0.0077 0.0077 0.0084 0.0126 0.0111 0.0119 0.0104 0.008
H2O 000000000000000000
MeOH 0 0.001 0.0015 0.0024 0.0024 0.0022 0.0021 0.0018 0.0017 0.0017 0.0015 0.0014 0.0013 0.0012 0.0012 0.0012 0.0012 0.0011
OME1 0.6647 0.6468 0.6512 0.6403 0.5948 0.529 0.4749 0.4089 0.355 0.3294 0.3188 0.309 0.3098 0.3141 0.3164 0.3179 0.321 0.3319
OME2 0 0.0168 0.01 0.0221 0.0671 0.1245 0.1666 0.2065 0.2282 0.231 0.2317 0.2271 0.2277 0.2304 0.2307 0.2322 0.2328 0.2345
OME3 0 0.0031 0.002 0.0043 0.0184 0.043 0.0683 0.1009 0.1301 0.1426 0.1486 0.1476 0.1482 0.1496 0.1493 0.1494 0.1488 0.1451
OME4 0 0.0026 0.0013 0.0031 0.0125 0.0279 0.0446 0.0674 0.0891 0.1003 0.1068 0.1072 0.1075 0.1081 0.1075 0.1069 0.1057 0.0995
OME5 0 0.0005 0.0003 0.0007 0.0035 0.0097 0.0179 0.0311 0.0462 0.0558 0.0618 0.0635 0.0639 0.064 0.0634 0.0625 0.0613 0.0556
OME6 0 0.0003 0.0002 0.0003 0.0014 0.0042 0.0083 0.0158 0.0259 0.0334 0.0385 0.0409 0.0413 0.0411 0.0406 0.0395 0.0385 0.0335
OME7 0 0.0001 0 0.0001 0.0007 0.002 0.0041 0.0081 0.0142 0.0192 0.023 0.0253 0.0256 0.0254 0.0249 0.0241 0.0231 0.0194
OME8 0 0 0 0 0.0003 0.0009 0.002 0.0041 0.0077 0.0109 0.0135 0.0154 0.0157 0.0155 0.0152 0.0145 0.0138 0.0111
OME9 0 0 0 0 0 0.0005 0.001 0.0022 0.0042 0.0063 0.0081 0.0094 0.0096 0.0096 0.0093 0.0088 0.0083 0.0064
OME10 0 0 0 0 0 0 0 0 0.0024 0.0036 0.0048 0.0057 0.0059 0.0059 0.0057 0.0054 0.0051 0.0038
OME11 0 0 0 0 0 0 0 0 0 0.0023 0.003 0.0037 0.0039 0.0039 0.0038 0.0036 0.0034 0.0024
TRI 0.3335 0.326 0.3312 0.3243 0.295 0.2512 0.2037 0.1463 0.0867 0.0529 0.0359 0.0314 0.0252 0.0238 0.0232 0.023 0.0231 0.0207
MEFO 0 0 0 0 0.0003 0.0006 0.001 0.0014 0.0022 0.0031 0.004 0.0046 0.006 0.0074 0.0086 0.011 0.014 0.0351
Table A.13:
OME synthesis from OME
1
-TRI over H-MFI 90 (conditions: OME
1
/TRI = 2
.
00
g g�1
, H-MFI 90/(OME
1
+TRI) = 1
.
0
wt
%, 60
◦C
,
8 bar, batch). The concentrations are presented in mass fractions and the duration in minutes.
t in min 0 0 1 5 10 15 20 30 45 60 75 90 120 150 180 240 300 1440
FA 0.0012 0.0018 0.002 0.0031 0.0035 0.0044 0.0073 0.0054 0.006 0.0064 0.007 0.007 0.0073 0.007 0.0067 0.007 0.0088 0.0066
H2O 000000000000000000
MeOH 0 0.0004 0.0009 0.0016 0.0016 0.0014 0.0013 0.0012 0.001 0.0009 0.0009 0.0009 0.0008 0.0008 0.0007 0.0007 0.0006 0.0005
OME1 0.6532 0.6551 0.6528 0.6404 0.6014 0.5553 0.5128 0.4494 0.3958 0.3711 0.355 0.3437 0.3299 0.3248 0.319 0.3191 0.3116 0.3419
OME2 0 0.0036 0.0043 0.0162 0.0585 0.1056 0.1448 0.1937 0.219 0.2259 0.2308 0.2297 0.2297 0.2317 0.2302 0.2305 0.2287 0.2343
OME3 0 0.0009 0.0007 0.0027 0.0127 0.0269 0.044 0.0775 0.1076 0.1219 0.132 0.1355 0.1414 0.1457 0.1463 0.1471 0.1472 0.1405
OME4 0 0.0008 0.0006 0.002 0.0074 0.0138 0.0215 0.0398 0.062 0.0751 0.0853 0.09 0.0975 0.1023 0.1037 0.1044 0.105 0.093
OME5 0 0.0002 0.0001 0.0002 0.0013 0.0033 0.0064 0.0153 0.0283 0.0371 0.0443 0.0483 0.0547 0.0585 0.0599 0.0606 0.0615 0.0501
OME6 0 0.0002 0 0.0001 0.0004 0.001 0.0022 0.0063 0.0138 0.0197 0.0249 0.028 0.0331 0.0363 0.0375 0.0381 0.0394 0.0292
OME7 0 0 0 0 0.0002 0.0004 0.0008 0.0025 0.0064 0.0101 0.0134 0.0156 0.0193 0.0215 0.0225 0.023 0.0243 0.0162
OME8 0 0 0 0 0 0 0.0002 0.001 0.003 0.0051 0.0072 0.0086 0.0111 0.0127 0.0134 0.0137 0.0148 0.0089
OME9 0 0 0 0 0 0 0 0.0004 0.0014 0.0026 0.0039 0.0048 0.0065 0.0076 0.0081 0.0083 0.0091 0.005
OME10 0 0 0 0 0 0 0 0 0.0007 0.0014 0.0021 0.0027 0.0038 0.0045 0.0049 0.005 0.0055 0.0028
OME11 0 0 0 0 0 0 0 0 0 0.0004 0.0012 0.0016 0.0024 0.0029 0.0031 0.0032 0.0036 0.0017
TRI 0.3456 0.3371 0.3385 0.3333 0.3123 0.2867 0.2571 0.2051 0.1513 0.1175 0.0934 0.0769 0.0546 0.0412 0.0333 0.0261 0.0248 0.0193
MEFO 0 0 0 0.0004 0.0008 0.0012 0.0017 0.0026 0.0037 0.0047 0.0057 0.0065 0.0079 0.0094 0.0106 0.0132 0.0151 0.0501
Table A.14:
OME synthesis from OME
1
-TRI over Nafon (conditions: OME
1
/TRI = 1
.
99
g g�1
, Nafon/(OME
1
+TRI) = 1
.
0
wt
%, 60
◦C
,8
bar
,
batch). The concentrations are presented in mass fractions and the duration in minutes.
t in min 0 0 1 5 10 15 20 30 45 60 75 90 120 150 180 240 300 1440
FA 0 0 0 0 0.0023 0.0035 0.0053 0.0073 0.0071 0.0078 0.01 0.0081 0.0091 0.0097 0.0086 0.01 0.0084 0.0098
H2O 0.0002 0.0002 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0.0005
MeOH 0 0.0004 0.0005 0.001 0.0012 0.0013 0.0015 0.0016 0.0019 0.0019 0.002 0.0021 0.0022 0.0022 0.0028 0.0028 0.0028 0.0024
OME1 0.6564 0.6627 0.6644 0.6396 0.5917 0.5464 0.5037 0.4424 0.3865 0.3549 0.3372 0.3271 0.3163 0.3091 0.314 0.3115 0.3098 0.294
OME2 0 0.0005 0.0038 0.0369 0.0873 0.1295 0.1635 0.205 0.229 0.2341 0.2337 0.2332 0.2308 0.2288 0.23 0.2291 0.2286 0.2304
OME3 0 0.0002 0.0003 0.0065 0.0211 0.0381 0.0561 0.0874 0.1179 0.1342 0.1415 0.1458 0.1484 0.1493 0.1492 0.1492 0.1493 0.1509
OME4 0 0.0002 0.0002 0.002 0.0072 0.0145 0.0236 0.043 0.0675 0.0848 0.0946 0.1006 0.1058 0.1082 0.1076 0.1079 0.1084 0.1094
OME5 0.0001 0 0 0.0004 0.0016 0.0039 0.0072 0.016 0.0301 0.0422 0.0502 0.0557 0.0612 0.0639 0.0633 0.0637 0.0643 0.0648
OME6 0 0 0 0.0001 0.0005 0.0012 0.0024 0.0063 0.0141 0.0221 0.0281 0.0326 0.0378 0.0406 0.0401 0.0407 0.0413 0.0416
OME7 0 0 0 0 0.0002 0.0004 0.0009 0.0025 0.0063 0.011 0.015 0.0182 0.0224 0.0248 0.0245 0.0251 0.0256 0.0258
OME8 0 0 0 0 0 0.0002 0.0003 0.001 0.0029 0.0054 0.0078 0.01 0.013 0.015 0.0148 0.0153 0.0158 0.0159
OME9 0 0 0 0 0 0 0 0.0005 0.0014 0.0028 0.0042 0.0057 0.0078 0.0093 0.0093 0.0098 0.0102 0.0101
OME10 0 0 0 0 0 0 0 0 0.0007 0.0014 0.0023 0.0032 0.0046 0.0056 0.0057 0.0061 0.0064 0.0063
OME11 0 0 0 0 0 0 0 0 0 0.0008 0.0013 0.0019 0.0029 0.0037 0.0038 0.0041 0.0043 0.0042
TRI 0.3433 0.3358 0.3308 0.3136 0.287 0.2611 0.2355 0.1868 0.1346 0.0966 0.0719 0.0558 0.0375 0.0296 0.026 0.0242 0.024 0.0251
MEFO 0 0 0 0 0 0 0 0 0 0 0 0 0.0002 0.0002 0.0003 0.0006 0.0009 0.0088
A.1.2 Thermal stability of the synthesis products
Table A.15 and A.16 list the bottom product composition of the distillations of the OME synthesis
products from the MeOH-pFA and OME
1
-TRI feed mixtures, respectively. Figure A.3 and A.4
illustrate the compositions of the bottom products of the distillations.
Table A.15:
Bottom product composition and conditions of the distillations of the OME synthesis
products from MeOH-pFA feed mixture. The concentrations are presented in mass
fractions.
A36 A46 Dowex H-BEA 25 H-MFI 90 Nafon
FA 0.38 0.41 0.36 0.39 0.33 0.4
H2O 0.1 0.12 0.1 0.1 0.11 0.1
MeOH 0.19 0.15 0.11 0.21 0.22 0.18
OME1 0.04 0.03 0.12 0.04 0.05 0.03
OME2 0.1 0.09 0.11 0.11 0.12 0.11
OME3 0.06 0.07 0.07 0.06 0.07 0.07
OME4 0.04 0.05 0.05 0.04 0.04 0.05
OME5 0.02 0.03 0.03 0.02 0.02 0.03
OME6 0.01 0.02 0.02 0.01 0.01 0.01
OME7 0.01 0.01 0.01 0.01 0.01 0.01
OME8 0 0.01 0.01 0 0 0
OME9 0 0 0 0 0 0
OME10 0 0 0 0 0 0
OME11 0 0 0 0 0 0
TRI 0.01 0.01 0.01 0 0 0
MEFO 0 0 0 0 0.01 0
minitial [g] 37.2 50.3 42 40.9 50.9 50.3
mDistillate [g] 3.8 8.1 3.2 2.5 6.6 5.5
mBottom [g] 32.9 41.9 38.3 38.3 43.8 44.3
TReboiler,max. [◦ C] 90 90 82 87 85 91
Figure A.3:
Bottom product composition and conditions of the distillations of the OME synthesis
products from MeOH-pFA feed mixture.
Table A.16:
Bottom product composition and conditions of the distillations of the OME synthesis
products from OME1-TRI feed mixture. The concentrations are presented in mass
fractions.
A36 A46 Dowex H-BEA 25 H-MFI 90 Nafon
FA 0.026 0.017 0.025 0 0 0.005
H2O 0 0 0 0 0 0
MeOH 0.006 0.01 0.01 0.001 0.002 0.002
OME1 0.113 0.035 0.351 0.172 0.154 0.084
OME2 0.289 0.129 0.214 0.296 0.313 0.3
OME3 0.189 0.176 0.133 0.186 0.197 0.199
OME4 0.136 0.186 0.093 0.128 0.132 0.145
OME5 0.081 0.14 0.052 0.073 0.072 0.087
OME6 0.051 0.106 0.032 0.045 0.042 0.057
OME7 0.032 0.075 0.02 0.027 0.024 0.036
OME8 0.02 0.05 0.013 0.016 0.013 0.022
OME9 0.012 0.034 0.011 0.009 0.008 0.014
OME10 0.008 0.023 0.011 0.005 0.004 0.009
OME11 0.005 0.016 0.015 0.003 0.003 0.006
TRI 0.031 0.015 0.03 0.027 0.025 0.034
MEFO 0.001 0 8E-04 0.011 0.012 0.001
minitial [g] 30.4 50.3 39 50.5 51.2 49.2
mDistillate [g] 7.1 9 7 11.5 10.3 11.8
mBottom [g] 23.1 40.7 30.5 38.4 40.3 37
TReboiler,max. [◦ C] 80 78 90 77 72 97
Figure A.4:
Bottom product composition and conditions of the distillations of the OME synthesis
products from OME1-TRI feed mixture.
A.2 Experimental demonstration of the main COMET process units
A.2.1 OME synthesis
Table A.17 lists the analytic results of the composition in the product barrels P1-P5. The
concentrations are presented in mass fractions.
Table A.17:
OME synthesis from OME
1
and concentrated FA(aq.) solution over A46 (conditions:
concentrated FA(aq.) with 85
�
89
wt
% FA, (concentrated FA(aq.) solution)/OME
1
=0
.
6
gg
�1
, A46/(OME
1
+concentrated FA(aq.) solution) = 0
.
34
gh g�1
, approx.
3
Lh
�1
, 90
◦C
, 10
bar
, fxed bed reactor). The concentrations are presented in
mass fractions.
P1-Exp P2-Exp P3-Exp P4-Exp P5-Exp
FA 0.2418 0.2434 0.1432 0.1526 0.1794
H2O 0.0366 0.0394 0.0394 0.0394 0.0394
MeOH 0.1476 0.1557 0.0996 0.1083 0.1111
OME1 0.3541 0.4141 0.2984 0.2808 0.2515
OME2 0.1068 0.0673 0.2006 0.1926 0.1818
OME3 0.0625 0.0413 0.1099 0.1116 0.1127
OME4 0.0293 0.0212 0.057 0.0597 0.0633
OME5 0.0131 0.0104 0.0291 0.0309 0.0341
OME6 0.0058 0.005 0.0146 0.0156 0.0179
TRI 0.0014 0.0011 0.0065 0.0065 0.0068
Tetroxane 0.0003 0.0003 0.0008 0.0008 0.0009
MEFO 0.0007 0.0009 0.0009 0.0011 0.0011
A.2.2 Synthesis product separation in CO-1
Table A.18 lists the analytic results of the composition of the distillate and bottom product
mixture of CO-1. The concentrations are presented in mass fractions.
Table A.18:
CO-1, OME synthesis product separation (conditions: 2
L h�1
, refux/distillate
= 0
.
5
�
2
s s�1
, distillate/feed = 81
wt
%, Montz 750 structured packing, 85
�
175
◦C
,
ambient pressure). The concentrations are presented in mass fractions.
Distillate Bottom
FA 0.196 0.0006
H2O 0.0471 0
MeOH 0.1973 0
OME1 0.2725 0
OME2 0.2456 0
OME3 0.0176 0.4458
OME4 0 0.2645
OME5 0 0.138
OME6 0 0.0714
TRI 0.0239 0
Tetroxane 0.0001 0.0052
MEFO 0 0
A.2.3 Reactive distillation in CO-2
Table A.19 lists the analytic results of the composition of the distillate and bottom product
mixture of CO-2. The concentrations are presented in mass fractions.
A.2.4 Product separation in CO-3
Table A.20 lists the analytic results of the composition of the distillate and bottom product
mixture of CO-3. The concentrations are presented in mass fractions.
Table A.19:
CO-2, Reactive distillation of the distillate product of CO-1 over A46 (conditions:
A46/(feed stream) = 0
.
35
gh g�1
,1
L h�1
, distillate/feed = 63
wt
%, Montz
750 structured packing, 45
�
104
◦C
, ambient pressure). The concentrations are
presented in mass fractions.
Feed Distillate Bottom
FA 0.2047 0 0.61
H2O 0.0486 0 0.38
MeOH 0.3031 0.0514 0.0035
OME1 0.1984 0.9486 0.0006
OME2 0.2202 0 0.0001
OME3 0.0158 0 0.0004
OME4 0 0 0.0005
OME5 0 0 0.0003
OME6 0 0 0.0002
TRI 0.0081 0 0.0006
Tetroxane 0 0 0.0038
MEFO 0.0011 0 0
Table A.20:
CO-3, product separation (conditions: 5
.
5
L h�1
, distillate/feed = 82
wt
%, Montz
750 structured packing, 100
�
210
◦C
, 200
mbar
). The concentrations are presented
in mass fractions.
Feed Distillate Bottom
FA 0.0006 0.0013 0
H2O 0 0.0001 0
MeOH 0 0.0003 0
OME1 0 0.0001 0
OME2 0 0.0002 0
OME3 0.4458 0.5184 0
OME4 0.2645 0.3402 0.0065
OME5 0.138 0.1296 0.257
OME6 0.0714 0.0021 0.3439
OME7 0.0368 0.0001 0.1942
OME8 0.0202 0 0.1052
OME9 0.0098 0 0.0537
OME10 0.0051 0 0.0269
OME11 0.0026 0 0.0124
TRI 0 0.0001 0
Tetroxan 0.0052 0.0075 0
MEFO 0 0 0
A.3 Process modelling and simulation
A.3.1 Pure component properties
Properties of the pure components used for the simulations are listed in Table A.21.
Table A.21: Pure component properties.
Component Parameters Reference
CO, CO2, FA, H2, H2O, MeOH, N2, O2, OME1
H2, MeOH, OME1, OME3-5
HF1
MG1
HFn, MGn, n > 1
OME2-10
OME2-5
cpig, g0, h0, ∆Vh, pc, pV, Tc, η, λ, ρ, σ
LHV
cpig, g0
, h0, pc, Tc
, η, λ, σ, ∆Vh, pV
cpig, g0
, h0, pc, Tc
, η, λ, σ, ∆Vh, pV
cp
ig, g0, h0, ∆ Vh, pc, Tc, η, λ, σ, pV
cpig, g0
, h0, pc, η, λ, σ, ∆Vh, pV, Tc
ρ
Aspen Database DB-PURE32
[109]
[108, 191]
[86, 108, 191, 192]
[108, 192]
[108, 127]
[98]
A.3.2 Thermodynamic model for mixtures
A UNIFAC based model for mixtures containing FA, MeOH and H
2
O was introduced by Maurer
et al. [126]. This model simultaneously considers the interactions between the components in
the liquid phase and the chemical reactions between FA and MeOH as well as FA and H
2
O
yielding the formation of poly-(oxymethylene) hemiformals and poly-(oxymethylene) glycols
following the reactions eqn. 2.1-2.4. Due to the fast kinetics of these reactions the assumption
that the equilibrium composition will be reached instantaneously shows good agreement with
the experimental results of vapor-liquid-equilibrium investigations. This model was further
developed in the following decades adding new components like TRI and OME
n
and adjusting
the interaction parameters to new experimental data. Recently Schmitz et al. [91] published
a new version of the model considering OME
n
. Bongartz et al. [108] implemented the model
version published by Kuhnert et al. [193] in Aspen Plus
®
and published the corresponding Aspen
Plus
®
fles. To include the chemical reactions of FA and MeOH as well as FA and H
2
O Bongartz
et al. [108] used the Chemistry section in Aspen Plus
®
which can be used to consider liquid
phase equilibrium reactions. The UNIFAC interaction parameters were slightly reformulated
to enable the implementation in Aspen Plus
®
, i.e. the temperature dependency was neglected,
instead the values at 300 K were considered. In this work, the Aspen Plus
®
model from Bongartz
et al. [108] were used and adjusted to consider the temperature dependency of the UNIFAC
interaction parameters. Therefore, the model to calculate the liquid phase activity coefcient
was adjusted to UNIFC-PSRK from the PSRK property model Gamma (GMUFPSRK) which
enables the consideration of temperature dependent interaction parameters and showed slightly
better results in the validation against experimental published results than the original model
from Bongartz et al. [108], see Table A.23. Furthermore, the UNIFAC parameters were adjusted
to the parameters published by Schmitz et al. [91]. The equation of the temperature dependent
UNIFAC interaction parameters for the sub-system H
2
O and CH
2
OH has the form of eqn. A.1,
where
ai,j
is the UNIFAC interaction parameter of the sub-system
i
, here H
2
O and
j
, here
CH
2
OH.
A
,
B
and
C
are the ftting parameters and
T
is the temperature in K. In Aspen Plus
®
the temperature dependency can be expressed according to eqn. A.2, therefore the equations
where adjusted and reftted. The results are presented in Table A.22 and Figure A.5.
ai,j
(T ) = A + B
T [K]
ai,j
(T ) = A + B · T [K] + C · T
2[K2]
(A.1)
(A.2)
Table A.22: Reft of UNIFAC interaction parameters.
A B C
a2,8 eqn. A.1, Literature 451.64 -114100 0
a2,8 eqn. A.2, Reft -521.15 2.7288 -0.0025
a8,2 eqn. A.1, Literature -1018.57 329900 0
a8,2 eqn. A.2, Reft 1794.1 -7.8899 0.0073
Figure A.5: Reft of UNIFAC interaction parameters.
Figure A.5 shows a good agreement between the reftted equation of the UNIFAC interaction
parameters and the equation from Schmitz et al. [91]. Deviations are in a far smaller range than
the values in the considered temperature range. The improvements regarding the description of
the phase behavior are presented in Table A.23.
A.3.3 Validation
For the validation of the implemented thermodynamic model describing the interactions in the
liquid and vapor phase several experimental VLE data from diferent literature sources were
used. The results are listed in Table A.23 and Figure A.6.
The validation was conducted using FLASH units in the simulation environment and running
sensitivity studies with cases containing the experimental data from the respective literature
sources. This procedure was chosen to enable the consideration of the formation of HF
n
and
MG
n
in sub-systems containing FA, MeOH and H
2
O. For the validation the overall composition
was calculated considering HF
n
and MG
n
as individual FA, MeOH and H
2
O molecules stoichio-
metrically. For a consistent procedure this approach was applied for sub-systems not containing
FA as well.
Table A.23:
Deviation of model predicted VLE data and experimental VLE data for four diferent models. The model Reference contains the model
predictions from the literature sources presenting the experimental data. This work contains the model predictions from the model
used for the process simulation in this work. [108] contains the model predictions from the model published by Bongartz et. al. [108].
[108]
∗
contains the model predictions from the model published by Bongartz et. al. [108] updated with the interaction parameters
from Schmitz et. al. [91], however still not considering the temperature dependency of the UNIFAC interaction parameters.
Sub-system Reference Data sets Model ∆Average ∆yF A ∆yMeOH ∆yH2O ∆yOM E1 ∆yOM E2 ∆T ∆p
FA/MeOH [194–196] 54 Reference 3.00 % 6.50 % 1.00 % - - - 0.40 % 4.10 %
This work 4.70 % 10.90 % 3.70 % 0.30 % 3.90 %
[108] 4.60 % 10.70 % 3.60 % 0.30 % 3.80 %
[108]∗ 4.60 % 10.70 % 3.60 % 0.30 % 3.80 %
FA/MeOH/H2O [126, 194, 197–199] 246 Reference 4.30 % 6.90 % 6.50 % 6.10 % - - 0.20 % 2.00 %
This work 4.20 % 7.40 % 5.80 % 5.50 % 0.20 % 2.20 %
[108] 5.60 % 10.70 % 7.80 % 6.20 % 0.30 % 3.00 %
[108]∗ 5.60 % 10.70 % 7.80 % 6.20 % 0.30 % 3.00 %
FA/MeOH/H2O/OME1 [191, 200] 45 Reference 21.10 % 26.40 % 19.70 % 8.90 % 67.70 % - 0.60 % 3.40 %
This work 14.70 % 26.10 % 19.10 % 10.90 % 55.40 % 0.80 % 5.40 %
[108] 15.20 % 27.90 % 22.20 % 8.80 % 55.00 % 0.60 % 6.70 %
[108]∗ 15.20 % 27.90 % 22.20 % 8.80 % 55.00 % 0.60 % 6.70 %
FA/MeOH/H2O/OME2 [91] 6 Reference - - - - - - -
This work 9.40 % 27.60 % 4.20 % 15.80 % 8.30 % 0.20 %
-
[108]∗ 7.00 % 26.00 % 4.00 % 3.50 % 8.40 % 0.20 %
FA/MeOH/OME1 [191] 17 Reference 10.00 % 16.30 % 7.30 % - 10.20 % - - 16.20 %
This work 11.60 % 24.90 % 4.70 % 10.10 % 18.10 %
[108] 14.90 % 38.00 % 3.60 % 23.80 % 23.00 % 0.60 %
[108] 11.10 % 23.90 % 4.40 % 9.50 % 17.50 %
[108]∗ 11.10 % 23.90 % 4.40 % 9.50 % 17.50 %
FA/H2O [86, 126, 191, 194, 201–204] 312 Reference 2.30 % 5.80 % - 1.10 % - - 0.40 % 1.90 %
This work 6.70 % 7.60 % 2.30 % 14.90 % 1.90 %
[108] 6.50 % 6.90 % 2.30 % 14.90 % 2.10 %
[108]∗ 6.50 % 6.90 % 2.30 % 14.90 % 2.10 %
FA/H2O/OME1 [191] 26 Reference 21.80 % 19.00 % - 13.90 % - - - 8.70 %
This work 19.60 % 14.80 % 15.20 % 5.90 %
[108] 20.20 % 16.40 % 14.90 % 6.30 %
[108]∗ 20.20 % 16.40 % 14.90 % 6.30 %
FA/H2O/OME2 [91] 6 Reference - - - - - - - -
This work 10.30 % 21.20 % 22.60 % 6.90 % 0.60 %
[108] 9.60 % 17.70 % 23.00 % 6.70 % 0.60 %
[108]∗ 9.60 % 17.70 % 23.00 % 6.70 % 0.60 %
MeOH/H2
O/OME1 [205] 23 Reference 5.00 % - 3.20 % 13.40 % 5.50 % - 0.20 % 2.60 %
This work 5.30 % 4.30 % 12.70 % 4.80 % 0.30 % 4.20 %
[108] 5.20 % 4.20 % 13.40 % 5.20 % 0.30 % 2.90 %
[108]∗ 5.20 % 4.20 % 13.40 % 5.20 % 0.30 % 2.90 %
MeOH/OME1 [191] 63 Reference 1.40 % - 2.00 % - 2.10 % - 0.10 % 1.40 %
This work 1.30 % 2.00 % 2.00 % 0.10 % 1.30 %
[108] 1.80 % 2.40 % 2.30 % 0.10 % 2.40 %
[108]∗ 1.80 % 2.40 % 2.30 % 0.10 % 2.40 %
MeOH/OME2 [206] 22 Reference - - - - - - - -
This work 1.60 % 1.90 % 4.60 % 0.10 %
[108] 4.00 % 4.20 % 11.60 % 0.40 %
[108]∗ 1.60 % 1.90 % 4.60 % 0.10 %
H2O/OME1 [191] 32 Reference 8.90 % - - 14.00 % 14.80 % - 0.20 % 6.60 %
This work 7.10 % 12.10 % 12.90 % 0.00 % 3.50 %
[108] 7.20 % 12.80 % 13.40 % 0.10 % 2.40 %
[108]∗ 7.20 % 12.80 % 13.40 % 0.10 % 2.40 %
OME1/OME2 [206] 21 Reference - - - - - - - -
This work 1.30 % 1.40 % 3.60 % 0.20 %
[108] 0.80 % 0.60 % 2.40 % 0.40 %
[108]∗ 0.80 % 0.60 % 2.40 % 0.40 %
Figure A.6:
Average deviation of VLE data from diferent sub-systems as presented in Table A.23
Following this approach, the knowledge of the feed composition, pressure and temperature level
is required. Since the experimental data usually only contain the composition of the liquid and
the vapor phase as well as temperature and pressure level but no information regarding the
mass distribution between liquid and vapor phase an assumption was made to defne the feed
composition. It was assumed that the feed composition is equal to the composition of the liquid
phase. To keep the resulting error small the vapor fraction in the fash unit was set to a small
value of 1
·
10
�4
to 1
·
10
�5
, depending on the sub-system. The choice of the vapor fraction was
based on the error between the resulting liquid phase composition to the liquid phase composition
from the experimental data which was generally smaller 0
.
1 %, only in a few cases the error
increases to 0.5 %.
The results in Table A.23 and Figure A.6 show that the model predictions agree well with
the experimental results. In addition to the deviations between the components, temperatures,
and pressure themselves, an average deviation is included enabling a fast comparison between
diferent models. The deviations of the diferent models are generally very close to each other.
A small improvement is visible for the model This work and [108]
∗
, which is mainly a result
of the updated model parameters from Schmitz et al. [91]. The behavior of most of the
sub-systems is described well by the model approach. However, systems containing the com-
ponents FA/MeOH/H
2
O/OME
1
, FA/MeOH/H
2
O/OME
2
, FA/MeOH/OME
1
, FA/H
2
O/OME
1
and FA/H
2
O/OME
2
show deviations of partly more than 20 %. This is also the case for the
Reference model, which presents the model predictions published together with experimental
data sets. Therefore, especially the predictions of the interactions between FA and OME show
potential for improvement with improved experimental data sets.
A.3.4 Mass balance and operation conditions of the process units of the COMET process
starting from H2 and CO2
A.3.4.1 MeOH sub-process
Figure A.7 illustrates a simplifed process fow diagram of the MeOH production based on H
2
and
CO
2
. Furthermore, it presents the stream numbering for the stream compositions and conditions
listed in Table A.24.
Figure A.7:
Simplifed process diagram for the production of MeOH from H
2
and CO
2
. CO,
distillation column; FL, phase separator; HE, heat exchanger; PC, compressor; R,
reactor.
Table A.25 to Table A.29 present the operation conditions of the main process units of the
MeOH sub-process including heat exchangers, a distillation column, reactor, phase separators
and compressors.
A.3.4.2 FA(aq.) sub-process
Figure A.8 illustrates a simplifed process fow diagram of the FA(aq.) production based on
MeOH and air. Furthermore, it presents the stream numbering for the stream compositions and
conditions listed in Table A.30.
Table A.31 to Table A.34 present the operation conditions of the main process units of the
FA(aq.) sub-process including heat exchangers, an absorber column, reactor and compressor.
A.3.4.3 FA concentration
Figure A.9 illustrates a simplifed process fow diagram of the FA concentration based on FA(aq.)
solution. Furthermore, it presents the stream numbering for the stream compositions and
conditions listed in Table A.35.
Table A.36 to Table A.38 present the operation conditions of the main process units of the
FA concentration sub-process including heat exchangers, a distillation column and evaporators.
Note that the operation conditions of the evaporators are diferent in practice with lower
operation pressure and higher operation temperature. However, the applied model is not suitable
Table A.24:
Stream table for the MeOH production based on H
2
and CO
2
. The stream numbering is presented in Figure A.7. The concentrations
are presented in mass fractions.
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19
T in ◦
C 59.5 25 240.1 250.1 40 47.7 40 62 60 131.9 47.7 60 60 84.1 58.2 58.2 58.2 59.9 99.5
p in bar 30 1 70 66.5 66.2 70 66.2 1.1 1 66.2 70 1 1.8 1.3 1 2.1 1.8 1.8 1
m in kg h�1 3142 22674 95791 95796 66852 69975 28943 33302 8189 3830 707 25113 11171 13942 8917 5288 3616 14787 5025
H2O 0 0 0.001 0.095 0.001 0.001 0.313 0.288 0.064 0 0.001 0.36 0.36 0.36 0.001 0.001 0.001 0.272 0.998
MeOH 0 0 0.006 0.174 0.009 0.008 0.555 0.592 0.45 0.007 0.008 0.639 0.639 0.639 0.998 0.999 0.999 0.727 0.002
H2 1 0 0.147 0.115 0.165 0.156 0 0 0.001 0.003 0.156 0 0 0 0 0 0 0 0
CO2 0 1 0.8 0.569 0.76 0.771 0.13 0.118 0.478 0.976 0.771 0.001 0.001 0.001 0.002 0 0 0.001 0
CO 0 0 0.046 0.047 0.066 0.063 0.002 0.002 0.006 0.014 0.063 0 0 0 0 0 0 0 0
Table A.25:
Operation conditions of the heat exchangers used for the MeOH production based
on H2 and CO2. Numbering is presented in Figure A.7.
HE-1 HE-2 HE-3 HE-4 HE-5 HE-6
Heat/cooling demand in MW 13.03 -13.52 2.14 3.21 1.69 -0.46
m in kg h�1 95791 95796 33302 13942 5288 5025
T1,in in ◦C 74.7 130.1 34.4 59.9 58.2 99.5
T1,out in ◦C 240.1 40 62 84.1 81 30
T2,in in ◦C 250.1 - - - - -
T2,out in ◦C 130.1 - - - - -
p1 in bar 70 66 1 1 2 1
p2 in bar 66 - - - - -
phase1,in gas gas/liquid gas/liquid liquid liquid liquid
phase1,out gas gas/liquid gas/liquid gas/liquid gas liquid
phase2,in gas - - - - -
phase2,out gas/liquid - - - - -
Table A.26:
Operation conditions of the distillation column used for the MeOH production based
on H2 and CO2. Numbering is presented in Figure A.7.
CO-1
Heat demand in MW 4.03
Cooling demand in MW -6.96
mFeed in kg h�1 13942
p in bar 1
TDist in ◦C 58.2
D:F in g g�1 0.64
TBott in ◦C 99.5
number of stages 28
refux:distillate in g g�1 1.5
Table A.27:
Operation conditions of the reactor used for the MeOH production based on H
2
and
CO2. Numbering is presented in Figure A.7.
R-MeOH
Cooling demand in MW -7.35
mFeed in kg h�1 95791
Tin in ◦C 240
Tout in ◦C 250
p in bar 70
Reactor type fxed bed reactor
Heat management isothermal
mCatalyst in kg 84112
GHSV in h�1 2639
Table A.28:
Operation conditions of the phase separators used for the MeOH production based
on H2 and CO2. Numbering is presented in Figure A.7.
FL-1 FL-2
Heat demand in MW 0 0
mFeed in kg h�1 95796 33302
p in bar 66 1
T in ◦C 40 62
Table A.29:
Operation conditions of the compressors used for the MeOH production based on
H2 and CO2. Numbering is presented in Figure A.7.
PC-1 PC-2 PC-3 PC-4
Power in MW 1.43 2.27 0.36 0.58
Cooling demand in MW 0 -1.67 0 -2.09
mFeed in kg h�1 3142 22674 70682 8189
pin in bar 30 1 66 1
pout in bar 70 70 70 66
Tin in ◦C 59.5 25 41.6 60
Tout in ◦C 173.1 132.1 47.7 131.9
Tintercooling in ◦C - 35 - 35
number of stages 1 4 1 4
Figure A.8:
Simplifed process diagram for the production of FA(aq.) from MeOH and air. CO,
distillation column; HE, heat exchanger; PC, compressor; R, reactor.
Table A.30:
Stream table for the FA(aq.) production from MeOH and air. The stream numbering
is presented in Figure A.8. The concentrations are presented in mass fractions.
1 2 3 4 5 6 7 8 9
T in ◦C 59.9 25 140 578.2 30.1 33.5 100.3 100.3 64.9
p in bar 1.8 1.8 1.4 1.3 1 1 1.8 1.8 1
m in kg h�1 14787 21881 64566 64566 1231 47287 19388 27899 18509
FA 0 0 0.001 0.142 0.184 0.003 0.003 0.003 0.501
H2O 0.272 0 0.077 0.15 0.795 0.033 0.033 0.033 0.491
MeOH 0.727 0 0.168 0.004 0.02 0.003 0.003 0.003 0.007
H2 0 0 0.004 0.007 0 0.01 0.01 0.01 0
CO2 0.001 0 0.025 0.042 0 0.058 0.058 0.058 0
CO 0 0 0.001 0.002 0 0.002 0.002 0.002 0
N2 0 0.71 0.587 0.587 0.001 0.801 0.801 0.801 0
O2 0 0.29 0.137 0.066 0 0.09 0.09 0.09 0
Table A.31:
Operation conditions of the heat exchangers used for the FA(aq.) production based
on MeOH and air. Numbering is presented in Figure A.8.
HE-1 HE-2 HE-3 HE-4 HE-5
Heat/cooling demand in MW 7.59 -10.56 -7.53 -3.23 -2.83
m in kg h�1 64566 64566 187143 113755 637748
Tin in ◦C 30.7 578.2 64.9 53.3 33.6
Tout in ◦C 140 160 30 30 30
p in bar 1 1 1 1 1
phasein gas/liquid gas liquid liquid liquid
phaseout gas gas liquid liquid liquid
Table A.32:
Operation conditions of the absorber column used for the FA(aq.) production based
on MeOH and air. Numbering is presented in Figure A.8.
CO-1
Heat demand in MW 0
Cooling demand in MW 0
mFeed in kg h�1 64566
p in bar 1
TDist in ◦C 33.5
D:F in g g�1 0.72
TBott in ◦C 64.9
number of stages 4
refux:distillate in g g�1 -
Table A.33:
Operation conditions of the reactor used for the FA(aq.) production based on MeOH
and air. Numbering is presented in Figure A.8.
R-FA
Cooling demand in MW 0
mFeed in kg h�1 64566
Tin in ◦C 140
Tout in ◦C 578
p in bar 1
Reactor type fxed bed reactor
Heat management adiabatic
mCatalyst in kg 2604
GHSV in h�1 15000
Table A.34:
Operation conditions of the compressor used for the FA(aq.) production based on
MeOH and air. Numbering is presented in Figure A.8.
PC-1
Power in MW 1.04
Cooling demand in MW 0
mFeed in kg h�1 47287
pin in bar 1
pout in bar 1.8
Tin in ◦C 33.5
Tout in ◦C 100.3
Tintercooling in ◦C -
number of stages 1
Table A.35:
Stream table for the FA concentration based on FA(aq.) solution. The stream
numbering is presented in Figure A.9. The concentrations are presented in mass
fractions.
1 2 3 4 5 6 7 8 9 10 11 12 13 14
T in ◦C 64.9 76.1 76.1 40 149.8 120.1 118.3 118.3 155.5 76.1 88.5 88.5 117.4 88.5
p in bar
m in kg h�1
1
18508
0.4
44273
0.4
14754
1.3
1231
5.5
13522
5.5
5652
1
2203
1
3449
5.5
7870
0.4
29520
0.5
45273
0.5
22316
1
15753
0.5
22957
FA 0.502 0.595 0.184 0.184 0.184 0.441 0.142 0.632 0 0.8 0.775 0.666 0.727 0.88
H2O 0.491 0.399 0.796 0.796 0.796 0.51 0.778 0.34 1 0.2 0.224 0.331 0.268 0.12
MeOH 0.007 0.007 0.02 0.02 0.02 0.048 0.08 0.028 0 0 0.002 0.003 0.005 0
Figure A.9:
Simplifed process diagram of the FA concentration based on FA(aq.) solution. CO,
distillation column; E, evaporator; HE, heat exchanger.
to accurately describe the behavior inside the evaporators. Therefore, ideal separator unit
operations were used to meet the mass balance and the operation conditions were adjusted to
meet the phase of the streams and the heat demand. A detailed description of an advanced
model for the simulation of the evaporators was published by Tönges et al. [153].
Table A.36:
Operation conditions of the heat exchangers used for the FA concentration based on
FA(aq.) solution. Numbering is presented in Figure A.9.
HE-1 HE-2 HE-3 HE-4 HE-5 HE-6 HE-7
Heat/cooling demand in MW -9.92 1.63 1.12 -1.36 -0.08 -1.37 -8.92
m in kg h�1 14754 13522 5652 2203 3449 7870 22316
Tin in ◦C 76.1 40 100.7 118.3 118.3 155.5 88.5
Tout in ◦C 40 149.8 118.3 30 90 30 86.2
p in bar 0.4 5.5 1 1 1 5.5 0.5
phasein gas liquid gas/liquid gas liquid liquid gas
phaseout liquid gas/liquid gas/liquid liquid liquid liquid liquid
A.3.4.4 OME3-5 sub-process
Figure A.10 illustrates a simplifed process fow diagram of the OME
3-5
sub-process based on
MeOH and concentrated FA(aq.) solution. Furthermore, it presents the stream numbering for
the stream compositions and conditions listed in Table A.39.
Table A.37:
Operation conditions of the distillation column used for the FA concentration based
on FA(aq.) solution. Numbering is presented in Figure A.9.
CO-1
Heat demand in MW 7.26
Cooling demand in MW -7.1
mFeed in kg h�1 13522
p in bar 5.5
TDist in ◦C 120.1
D:F in gg
�1 0.42
TBott in ◦C 155.5
number of stages 32
refux:distillate in gg
�1 1.2
Table A.38:
Operation conditions of the evaporators used for the FA concentration based on
FA(aq.) solution. Numbering is presented in Figure A.9.
E-1 E-2 E-3
Heat demand in MW 9.31 8.89 0
mFeed in kg h�1 44273 45273 5652
p in bar 0.4 0.5 1
T in ◦C 76.1 88.5 118.3
Figure A.10:
Simplifed process diagram for the production of OME
3-5
from MeOH and concen-
trated FA(aq.) solution. CO, distillation column; HE, heat exchanger; R, reactor.
Table A.39:
Stream table for the production of OME
3-5
from MeOH and concentrated FA(aq.)
solution. The stream numbering is presented in Figure A.10. The concentrations
are presented in mass fractions.
1 2 3 4 5 6 7 8 9 10 11
T in ◦C 90.4 90 90 81.5 81 69.5 41.5 117.4 200.5 86.6 194.9
p in bar 10.3 10 10.1 1.8 1.8 1 1 1 1.8 0.1 0.1
m in kg h�1 22957 66666 66666 51796 5288 57083 41330 15753 14871 12490 2380
FA 0.88 0.303 0.186 0.239 0 0.217 0 0.727 0 0 0
H2O 0.12 0.042 0.022 0.028 0 0.026 0.002 0.268 0 0 0
MeOH 0 0.028 0.1 0.129 1 0.21 0.045 0.005 0 0 0
OME1 0 0.591 0.276 0.356 0 0.323 0.953 0 0 0 0
OME2 0 0 0.179 0.23 0 0.209 0 0 0 0 0
OME3 0 0 0.107 0.017 0 0.015 0 0 0.419 0.499 0
OME4 0 0 0.061 0 0 0 0 0 0.271 0.323 0
OME5 0 0 0.033 0 0 0 0 0 0.149 0.177 0
OME6 0 0.018 0.018 0 0 0 0 0 0.08 0 0.496
OME7 0 0.009 0.009 0 0 0 0 0 0.042 0 0.262
OME8 0 0.005 0.005 0 0 0 0 0 0.022 0 0.136
OME9 0 0.002 0.002 0 0 0 0 0 0.011 0 0.07
OME10 0 0.001 0.001 0 0 0 0 0 0.006 0 0.036
Table A.40 to Table A.42 present the operation conditions of the main process units of the
OME
3-5
sub-process including heat exchangers, distillation columns and a reactor. Note that
the WHSV of the OME reactor is overestimated and much lower in practice. The complexity is
described in section 5.2.1.
Table A.40:
Operation conditions of the heat exchangers used for the production of OME
3-5
from
MeOH and concentrated FA(aq.) solution. Numbering is presented in Figure A.10.
HE-1 HE-2 HE-3 HE-4 HE-5
Heat/cooling demand in MW 1.52 5.01 0.12 0.39 -0.32
m in kg h�1 66666 66666 57083 14871 12490
Tin in ◦C 58.6 90 76.4 120.2 86.6
Tout in ◦C 90 130 69.5 130 30
p in bar 10 2 1 0.1 1
phasein liquid liquid gas/liquid gas/liquid liquid
phaseout liquid gas/liquid gas/liquid gas/liquid liquid
A.3.4.5 Combustion
Figure A.11 illustrates a simplifed process fow diagram for the combustion of the purge streams.
Furthermore, it presents the stream numbering for the stream compositions and conditions listed
in Table A.43.
Table A.41:
Operation conditions of the distillation columns used for the production of OME
3-5
from MeOH and concentrated FA(aq.) solution. Numbering is presented in Figure
A.10.
CO-1 CO-2 CO-3
Heat demand in MW 12.69 4.47 0.47
Cooling demand in MW -17.17 -7.9 -1.67
mFeed in kg h�1 66666 57083 14871
p in bar 1.8 1 0.07
TDist in ◦C 81.5 41.5 86.6
D:F in g g�1 0.78 0.72 0.84
TBott in ◦C 200.5 117.4 194.9
number of stages 56 40 30
refux:distillate in g g�1 0.5 0.7 0.3
Table A.42:
Operation conditions of the reactor used for the production of OME
3-5
from MeOH
and concentrated FA(aq.) solution. Numbering is presented in Figure A.10.
R-OME
Heat demand in MW 0.49
mFeed in kg h�1 66666
Tin in ◦C 90
Tout in ◦C 90
p in bar 10
Reactor type fxed bed reactor
Heat management isothermal
mCatalyst in kg 952
WHSV in h�1 70
Figure A.11:
Simplifed process diagram for the combustion of the purge streams. HE, heat
exchanger; PC, compressor; R, reactor.
Table A.43:
Stream table for the combustion of the purge streams. The stream numbering is
presented in Figure A.11. The concentrations are presented in mass fractions.
123
T in ◦C 25 54.6 98.5
p in bar 1 1 1
m in kg h�1 34730 54838 54838
FA 0 0.001 0
H2O 0 0.012 0.063
MeOH 0 0.001 0
H2 0 0.005 0
CO2 0 0.031 0.036
CO 0 0.002 0
N2 0.742 0.753 0.753
O2 0.258 0.195 0.147
Table A.44 to Table A.46 present the operation conditions of the main process units for the
combustion of the purge streams including heat exchangers, a combustion chamber and a
compressor.
Table A.44:
Operation conditions of the heat exchangers used for the combustion of the purge
streams. Numbering is presented in Figure A.11.
HE-1 HE-2 HE-3
Heat/cooling demand in MW -9.68 -1.16 -1.32
m in kg h�1 54838 54838 54838
Tin in ◦C 772.9 230 160
Tout in ◦C 230 160 98.5
p in bar 2 1.7 1.3
phasein gas gas gas
phaseout gas gas gas
A.3.5 Comparison to alternative OME3-5 production processes
The key assumptions for the overall energy efciency of various OME
3-5
production processes
including the production of H
2
via H
2
O electrolysis and various CO
2
capture techniques are
summarized in Table A.47. The assumptions were considered from Held et al. [109].
Table A.45:
Operation conditions of the reactor used for the combustion of the purge streams.
Numbering is presented in Figure A.11.
R-Combustion
Heat demand in MW 0
mFeed in kg h�1 54838
Tin in ◦C 151
Tout in ◦C 773
p in bar 2
Reactor type Combustion chamber
Heat management steam generation
Table A.46:
Operation conditions of the compressor used for the combustion of the purge streams.
Numbering is presented in Figure A.11.
PC-1
Power in MW 1.61
Cooling demand in MW 0
mFeed in kg h�1 54838
pin in bar 1
pout in bar 2.1
Tin in ◦C 54.6
Tout in ◦C 150.7
Tintercooling in ◦C -
number of stages 1
Table A.47:
Energy demand for H
2
O electrolysis and various CO
2
capture techniques [109]. CPS,
CO2 from point sources; PCC, postcombustion capture; DAC, direct air capture.
H2O electrolysis CPS PCC DAC
Electricity demand in MJ kg�1 product 200.2 0 0 0.9
Heat demand in MJ kg�1 product 0 0 3.33 6.3
