RESEARCH TRIANGLE INSTITUTE
July 1988
Phase I — Final Report
Catalytic Dehydration
of Methanol
Prepared by
Don R. vanderVaart
Research Triangle Institute
Research Triangle Park, NC 27709
Prepared for
Robert M. Heavenrich, Project Officer
U.S. Environmental Protection Agency
Ann Arbor, Ml 48105
EPA Contract No. 68-03-3527
RTI Project No. 4002
POST OFFICE BOX 12194 RESEARCH TRIANGLE PARK, NORTH CAROLINA 27 709-2194
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CONTENTS
CONTENTS
Section
Page
Figures iii
Tables iv
Abstract v
1 Introduction 1-1
2 Candidate Catalysts.... 2-1
Alumina 2-1
Zeolite 2-6
Metal Oxides 2-6
Ion Exchange Resins 2-6
3 Experimental 3-1
Experimental Considerations 3-2
Experimental Procedure 3-5
4 Results 4-1
5 Discussion 5-1
6 Phase II Plan of Work 6-1
Dissociation Reaction .- 6-7
Process Energetics 6-10
7 References 7-1
8 Summary and Future Work 8-1
Appendix
A Solicitation for Commercial Methanol Dehydration Catalysts A-l
B Calculation of Adiabatic Temperature Rise in a Tubular Reactor B-l
C Calculation of Methanol Required to Start an Engine C-l
D Estimation of Global Kinetics of Methanol Dehydration D-l
E Calculation of the Vapor-Liquid Equilibrium (VLE) Data for a
Two Component Mixture of DME and Methanol E-l
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FIGURES
FIGURES
Number Page
1 The effect of spare time on product distribution (T = 371 °C, P = 1 atm) 1-2
2 Microreactor catalyst testing facility 3-6
3 10-Port sampling valve arrangement 3-7
4 MeOH conversion, AL-5207 4-2
5 MeOH conversion, AL-5307 E 1/16 4-2
6 MeOH conversion, AL-5407 4-3
7 MeOH conversion, AL-3945 4-3
8 MeOH conversion, AL-3996 R 4-4
9 MeOH conversion, CS 331-1 4-4
10 MeOH conversion, CS331-4 4-5
11 MeOH conversion, LZ-20 4-5
12 MeOH conversion, LZ-Y-72 4-6
13 MeOH conversion, LZ-M 8 4-6
14 MeOH conversion, Z6-06-02 4-7
15 MeOH conversion. Z6-06-02,B6980 4-7
16 MeOH conversion, MCG-7 4.8
17 MeOH conversion, MCG-8 4-8
18 MeOH conversion, ZR-0304T 4-9
19 MeOH conversion, TI-0720 4-9
20 MeOH conversion, T-312 4-10
21 MeOH conversion, T-314 4-10
22 MeOH conversion, T-317 4-11
23 MeOH conversion, T-1502A 4-11
24 MeOH conversion, T-1502B 4-12
25 MeOH conversion. 7913-S K-306 4-12
26 MeOH & DME outputs, AL-5207 4-13
27 MeOH & DME outputs, AL-5407 E 1/16 4-1-3
28 MeOH & DME outputs, LZ-20 4-14
29 MeOH & DME outputs, M-8 4-14
30 MeOH & DME outputs, Z6-06-02 1/16 4-15
31 Effect of Tc on the temperature profile of a PFTR 6-4
32 Relationship between two dimensionless parameters 6-5
33 MeOH conversion, dissociation catalyst 6-8
34 MeOH conversion, dissociation and AL 39996R 6-8
35 MeOH conversion, dissociation in DME 6-9
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TABLES
TABLES
Number Page
1 Advantages and Disadvantages of Candidate Catalyst Families 2-2
2 Catalyst Samples Received and Tested 2-4
3 Harshaw/Fitrol Alumina Catalysts 2-5
4 United Catalyst Alumina Catalysts 2-5
5 Union Carbide Zeolite Catalysts 2-7
6 United Catalyst/Zeochem Zeolite Catalysts 2-7
7 Harshaw/Filtrol Metal Oxide Catalysts 2-8
8 United Catalysts Metal Oxide Catalysts 2-8
9 Ranking of All Candidate Catalysts Tested According to T30 4-16
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ABSTRACT
ABSTRACT
The results of Phase I of EPA Contract No. 68-03-3527 are presented. It is
apparent that several commercially available catalysts are suitable for the
methanol dehydration reaction forming dimethyl ether. Problems in
developing a prototype dehydration reactor suitable for mounting under the
hood of an automobile are presented. In particular, the exothermicity of the
methanol dehydration reaction can result in a runaway temperature excursion
within the reactor. This can lead to the loss of product selectivity by the
formation of higher hydrocarbons with the concomitant formation of coke,
which can lead to catalyst deactivation. Although more sophisticated reactor
designs can be used to minimize the likelihood of these excursions, the reaction
system is quite sensitive to slight changes in operating conditions.
It is proposed, in Phase II, to study the kinetic rate of both the methanol
dehyrdation reaction for a limited group of dehydration catalysts as well as the
methanol dissociation reaction over available dissociation catalysts. With this
information, a relatively simple theoretical analysis is feasible by which a
dehydration/dissociation catalyst mixture reactor could be designed. The
dissociation reaction pathway would be provided only as a heat sink to prevent
reactor runaway.
Finally, it is shown that the process energetics for such a reactor would require
the condensation of the product stream in order to allow operation without a
significant amount of external power requirements. A simple phase
equilibrium analysis is presented which shows that the condensed product
stream is suitable for cold-starting an internal combustion in a mechanism
completely analogous to winterized gasoline.
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1. INTRODUCTION
SECTION 1
INTRODUCTION
Methanol can be used as a motor fuel with considerable environmental
advantages over gasoline. Its use in various mixtures with gasoline is currently
being tested by both the Ford Motor Company and ARCO. Using 100% meth-
anol rather than blending it with a winterized gasoline, while representing a tru-
ly alternative fuel with excellent emission characteristics, does present certain
problems. In particular, the high heat of vaporization of methanol (more than
seven times that of gasoline on a per Btu heating value basis) combined with
the higher fuel to air ratio required for combustion (approximately twice that of
gasoline) causes severe cold starting problems. To this end the U.S. Environ-
mental Protection Agency (EPA) has suggested that under cold starting
conditions an amount of methanol fuel could be catalytically dehydrated to
form dimethyl ether (DME) via the following reaction:
2CH3OH «- CH3OCH3 + H20 . (1)
DME exhibits favorable combustion characteristics and is satisfactorily
volatile.
The catalytic dehydration of methanol is a well-studied reaction. It forms the
first step in Mobil's patented Methanol to Gasoline (MTG) process as well as
being an integral part of other chemical production processes. From the litera-
ture relative to reaction (1) two conclusions can be drawn:
• Catalytic dehydration of methanol vapor to dimethyl ether, is an acid
catalyzed reaction. Lewis acid sites and almost certainly fairly weak Lewis
base sites are necessary on the surface of any successful catalyst. Strong
Bronsted acidity promotes the further reaction of DME to olefins and
paraffins. _
• While most research on this reaction has focused on various conventional
zeolites, other catalysts have been shown to be successful. These include
metal salts, alumina, silica alumina, alumino phosphate molecular sieves,
ion exchange resins, and mixed metal oxides.
As an example of the work published in the development of Mobil's MTG
process, Figure 1 shows the effect of the space time (liquid hourly space velocity
(LHSV)'1) (defined as the volumetric flowrate of liquid methanol divided by
RTI 1-1
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Paraffins
(and Cg Olefins)
3)
O
o
c
o
Figure 1. The effect of space time on product distribution (T = 371 °C, P = 1 atm) [1].
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1. INTRODUCTION
the volume of catalyst) on product distribution over a zeolite catalyst. [1]
Figure 1 also illustrates the series reactions that can occur in this system. That
is
A*B + C (2)
where A is methanol, B is dimethyl ether, and C constitutes a variety of hydro-
carbon products which make up the gasoline fraction produced by the MTG
process. Obviously, it is this latter product fraction which Mobil attempts to
maximize. The formation of these hydrocarbons (aromatics, olefins, and some
paraffins) has associated with it, however, the formation of carbonaceous
deposits on the catalysts (i.e., coke). This byproduct has the effect of
deactivating the catalyst over time, necessitating its regeneration. The goal of
the present work is not to produce gasoline in situ but rather to provide a
volatile product suitable for starting the vehicle in cold weather. Dimethyl ether
formed as an intermediate in reaction (2) meets this objective without the
associated deactivation of the catalyst. In fact, Mobil has remarked [2] that
their first stage unit, which produces DME at equilibrium conversion, has yet to
experience catalyst deactivation at their plant in New Zealand, while their
second stage DME-to-gasoline reactor requires continuous regeneration. The
work presented herein attempts to produce a methanol/DME mixture suitable
for starting an engine. .
Previous work [3] indicates that this mixture should be roughly 10 to 15 percent
(by volume) DME in methanol. In terms of Figure 1, then, a reactor should
operate with a space time (1/LHSV) of roughly 10~3 hr under the conditions
studied in that work, (i.e., 371° C 1 atm).
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2. CANDIDATE CATALYSTS
SECTION 2
CANDIDATE CATALYSTS
Alumina
A discussion of the catalytic chemistry involved in the dehydration of methanol
to dimethyl ether is beyond the scope of this report. Several investigations of
this reaction by various researchers [1,4,5] have been reported. In the particular
application of methanol dehydration to facilitating the cold starting of an
automobile engine, three criteria that the optimal catalyst should meet are:
• the catalyst must be sufficiently active to produce a product stream capable
of satisfying the cold starting requirements of an engine (roughly 15%
DME in methanol).
• the catalyst must be stable under temperature excursions.
• the catalyst must be economical when normalized by its expected lifetime.
Deactivation will likely occur due either to coking reactions caused by
further dehydration of DME to unsaturated hydrocarbons (reaction (2)) or
to high temperature sintering.
A list of various catalyst families suitable for the dehydration reaction along
with some of their advantages and disadvantages is presented in Table 1.
Upon award of this contract, a number of commercial catalyst vendors were
contacted to obtain catalysts suitable for the methanol dehydration reaction.
An example of the request is given in Appendix A. -Of the catalyst families
listed in Table 1, only four were represented by samples received in response to
this solicitation. These are presented in Table 2 categorized by family. A brief
description of these catalysts, including the limited information provided by
each vendor, is presented below.
Aluminum oxide (A12O3) exhibits significant acid properties suitable for metha-
nol dehydration reaction. These acid properties can be varied somewhat by the
calcining procedure used in their manufacturing so that a large variety of
aluminas exist. Harshaw/Filtrol provided a number of alumina catalysts as
listed in Table 2. A summary of the composition of each of these catalysts is
presented in Table 3. The properties of the United Catalyst alumina catalysts
are presented in Table 4. Note that the surface areas do vary somewhat.
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2. CANDIDATE CATALYSTS
Table 1. Advantages and Disadvantages of Candidate Catalyst Families
Catalyst
Advantages
Disadvantages
Zeolites
aluminosilicates
aluminophosphates
Silica-aluminas
Aluminas
_ Metal salts
phosphates
sulfates
Much research has been done Possible selectivity
on this catalyst; methanol problems since DME
dehydration is fairly well readily react further
understood on these catalysts
Probably has the right mix
of acid-base site strengths
Well-studied catalysts for
this reaction
Bronsted/Lewis acid site
ratio is controllable by
calcining conditions
Low cost catalyst
Acid strength distribution
is controllable by calcin-
ing temperature
Bronsted/Lewis acid/base
character is controllable
by dehydroxylation
May have surface acidity
comparable to alumina
Bronsted/Lewis acid-base
sites are controllable by
calcination
Relatively new
materials; methanol
dehydration has not
been well studied
High density of very
strong acid sites may
lead to low
selectivity and/or
coking
Selectivity may be
poor; alumina is very
active for olefin
formation from
alcohol dehydra-
tion
Small amounts of
impurities can affect
the acid site strength
and thus the activity/
selectivity
Selectivity may be
poor due to strength
of acid sites
Inherently strong acid
sites may lead to
coking
(continued)
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2. CANDIDATE CATALYSTS
Table 1 (continued)
Catalyst Advantages Disadvantages
Ion exchange resins High activity at low tempera- Poor mass transfer
tures in some resins due
to small pore sizes
Acidity is controllable by Deactivation may
synthesis conditions occur if temperature
excursions above
~150° C occur
Mixed metal oxides Acid site strength con- Selectivity may be
trollable by synthesis poor since some
conditions mixed oxides also
promote dehydroge-
nation and olefin
formation.
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2. CANDIDATE CATALYSTS
Table 2. Catalyst Samples Received and Tested
Aluminas
Zeolites
Metal oxides
Harshaw AL-5207
AL-5307
AL-5407
AL-3945
AL-3996R
United CS 331-1
CS 331-4
Union Carbide LZ-20
LZ-Y-72
LZ-M8
United Z6-06-02
Mobil MCG-7
MCG-8
Zeochem Z6-06-02, B6980
Harshaw ZR-0304T
TI-0720
United T-312
T-314
T-317
T-1502A
T-1502B
7913-S, K-306
Ion Exchange Resins
Rohm & Haas
XE-386
6-928 Not tested
6-9852
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2. CANDIDATE CATALYSTS
Table 3. Harshaw/Filtrol Alumina Catalysts
Catalysts
AL-5207
AL-5307
AL-5407
AL-3945
AL-3996R
Surface Area
(m2/g)
200
200
190
250
200
Fe203
0.1
0.1
0.1
< 0.01
0.07
Composition (wt%)
w/Balance AI203
SiO2 Na2O
0.2
0.2
0.2
<0.01
0.01
0.08
0.08
0.08
<0.01
0.06
Table 4. United Catalyst Alumina Catalysts
Catalysts
CS 331-1
CS 331-4
Surface Area
(m2/g)
200-300
200-300
Fe203
<0.20
<0.20
Composition (wt%)
w/Balance AI2O3
Na Cl S
<0.05 <0.02 <0.05
<0.05 <0.02 <0.05
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2. CANDIDATE CATALYSTS
Zeolite
Metal oxides
The term zeolite refers to crystalline aluminosilicate compounds which have
well-defined pore structures and geometries. They also contain an exchange-
able metal cation and considerable water of hydration as synthesized. By
varying the cation and the amount of water present in the structure during
manufacturing, the acid/base properties can be somewhat controlled.
Table 2 lists the zeolites received. Their properties are given in Tables 5 and 6.
It is possible that the Mobil samples (MCG-7 and 8) are variations of their
ZSM-5 zeolite which is used in their second stage MTG reactor. No infor-
mation was provided by Mobil in this regard, however, so that positive
identification is impossible.
Metal oxides are of particular interest for this application because the acidity
and basicity of the catalyst can be easily altered through changes in catalyst
formulation. A number of metal oxide catalysts on different supports were
received as shown in Table 2. Their properties are shown in Tables 7 and 8.
Ion Exchange Resins
Ion exchange resins have shown low temperature (30° -150° C) activity for
alcohol dehydration [4]. These catalysts are extremely sensitive to both high
temperature (> 150° C) and hydration effects, so no testing was undertaken at
this time. In addition, the ion exchange resins cannot be formed into monoliths
thereby disqualifying them from consideration in this automotive application.
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2. CANDIDATE CATALYSTS
Table 5. Union Carbide Zeolite Catalysts
Catalysts
LZ-20
LZ-M8
LZ-Y-72
Composition (wt%)
AI2O3 Si02
<40 <80
-
<40 <80
Na20
<5
-
<5
NHL Remarks
< 20 Ammonium
Exchanged
-
Table 6. United Catalyst/Zeochem Zeolite Catalysts
Catalysts
Surface Area
(m2/g)
AI203
Composition (wt%)
Na2OAI2O3x
Si02 x H20
MO
United Z6-06-02
Zeochem Z6-06-02
B6980
375
25-35
25-35
65-75
65-75
2.4
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2. CANDIDATE CATALYSTS
Table 7. Harshaw/Filtrol Metal Oxide Catalysts
Catalysts
ZR-0304T
0720
Surface Area
m*/g
52
205
Composition (wt%)
ZrO2 Ti02 AI2O3 Si02
980 __,.. 2.0 0.1
100
Table 8. United Catalysts Metal Oxide Catalysts
Catalysts
T-312
T-314
T-317
T-1502A
T-1502B
K-306
Surface Area
m2/g
183
. 190
187
124
152
250
Ni
8-10
8-10
15
27
Si02
71.7
Composition (wt%)
Cu Cr Ti
10
10-12
AI203
12.5
1.6
0.1-0.2
Fe2O3 CaO
5.2 2.7
Balance
A1203
A1203
A1203
SiO2
SiO2
MgO
3.6
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3. EXPERIMENTAL
SECTION 3
EXPERIMENTAL
The primary objective of Phase I was to screen and rank a variety of commer-
cially available catalysts with the potential of effecting the dehydration of meth-
anol to dimethyl ether (DME). The minimum requirement of a catalytic dehy-
dration system is to deliver a gas stream of roughly 15 percent DME in metha-
nol to the automobile fuel system under cold start conditions. This
corresponds to 30% conversion of methanol by reaction (1). Given that
objective and the constraints of available space onboard an automobile, the
catalyst must exhibit sufficient activity to produce the required stream in a reac-
tor of feasible size. In addition, to reduce high temperature deactivation, the
catalyst must convert approximately 30 percent of the methanol feed at as low a
temperature as possible. The rate at which methanol will dehydrate is a
function of the reactor temperature and catalyst. Once these are fixed, the ex-
tent of conversion of methanol to DME depends on the product of this rate
times the time which the methanol spends in contact with the catalyst. This
contact time is conveniently given by the liquid hourly space velocity (LHSV).
The LHSV is defined as:
T ncv - the volume of liquid methanol/time /«->
LHb ~ the volume of catalyst^ }
The overall conversion is then given by :
(conversion of methanol in moles) = r^nov (4)
where:
r = overall reaction rate of methanol dehydration in moles converted per
unit time
The reaction rate (r) is normally given in the form of:
r = Ae E/RT (C)n (5)
where:
A = pre-exponential factor which is approximately constant
E = activation energy of the reaction
R = gas constant
T = reaction temperature
C = concentration of the reactants
n = order of reaction
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3. EXPERIMENTAL
For a given methanol feedstream and catalyst at constant temperature, r can be
considered constant. Thus, the conversion of methanol can be increased by
decreasing LHSV by Equation (4). Physically, this means that for a given
reaction rate, a lower flow rate will permit a higher conversion of methanol.
Alternatively, the reaction rate for a given catalyst can be increased dramatical-
ly by increasing the reactor temperature, T. The activity of a catalyst is
manifested by the activation energy, E. By Equation (5), a lower E has the
effect of increasing the reaction rate at a given temperature. Hence, a catalyst
that is more "active" exhibits a lower activation and therefore a higher reaction
rate for a given temperature, T. Consequently, a more active catalyst will in-
crease conversion of methanol for a given reactor size and capacity (LHSV).
Finally, the reaction order (n) can be considered to be equal to one for the het-
erogeneously catalyzed methanol dehydration reaction as a first approximation.
In the application of onboard methanol dehydration, the total volume of the
proposed catalytic reactor must be small. In addition, the reaction is intended
to operate under cold start conditions so that a minimum reaction temperature
is sought. It should be noted that for a given catalyst this presents a tradeoff in
that the lower the reaction temperature, the higher the reactor volume must be
to maintain a given level of methanol conversion. The objective of Phase I,
then, is simply to identify those catalysts that exhibit the lowest activation ener-
gy for the methanol dehydration reaction. These will, in turn, provide the best
level of conversion for the minimum reaction temperature and reaction
volume.
Experimental Considerations
The rate of a heterogeneously catalyzed reaction of gas phase reactants on the
surface of the solid catalyst is determined by the slowest of the following steps:
• External mass transfer - the rate at which either the gas phase reactants reach
or the products leave the surface of the catalyst. This is a function of both
the bulk motion of the gas around the catalyst particle as well as the
diffusion properties of the reactants and products under the experimental
conditions.
• Intraparticle mass transfer - once the gas phase reactants have reached the
catalyst pellet, they must find their way to an active site on the surface.
Similarly, after reaction, the products must diffuse away from the active site.
• Surface reaction rate - Once the gas phase reactant has reached an active site
on the catalyst, it must react to form the reaction products. The inherent
activation energy describes the rate of this reaction and varies from catalyst
to catalyst.
The relative contribution of each of these steps is most easily visualized by con-
sidering a varying reactor temperature. Obviously, at very low temperatures.
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3. EXPERIMENTAL
the reaction rate on the surface is extremely slow and therefore rate controlling.
The conversion of the gas phase reactants is limited by the reaction rate on the
surface of the catalyst. As the temperature is increased, however, this reaction
rate increases (Equation [5]) to the point at which it exceeds the rate at which
the reactants reach the surface. Mass transfer is said to be limiting the reaction
rate. This resistance could be the rate of reactant transfer from the bulk gas
stream to the catalyst pellet (external mass transfer). Most catalysts are porous
so that the mass transfer into the pellet pores to the actual surface could also be
limiting the overall reaction rate (internal mass transfer). Usually, however, the
surface area of the catalyst is large and the particles are small, in which case
external mass transfer from the bulk gas stream to the catalyst particle is rate
limiting.
In any event, practical systems utilize high reactor temperatures to ensure that
the surface chemical kinetics do not limit the production rate. The overall mass
transfer of the reactants to the catalyst particle is, therefore, oftentimes the rate
limiting step in these heterogeneous systems.
For the purposes of the Phase I screening tests, it was determined to fix as
many of the experimental factors as possible so that the evaluation of each cata-
lyst would be under uniform conditions. This need for constancy was effected
by a number of experimental considerations each of which will be treated sepa-
rately.
The LHSV as defined by Equation (3) was kept constant for each catalyst test-
ed. This was done by using the same volume of each catalyst for each test while
maintaining a constant methanol volumetric flow rate. In particular, a 0.4 cm
I.D. glass tube was packed with catalyst to a length of 4 cm. The volumetric
flow rate was constant at 200 cm3/min of 1,000 ppm methanol in nitrogen. By
Equation (3), this equals a LHSV of 6.08 x 10"2 hr"1 (see Section 6).
Within the catalyst bed, the amount of time that the methanol feed spends in
contact with the candidate catalyst should be uniform over the cross-section of
the packed reactor. In this study, the candidate catalysts were crushed and
sieved to produce a 60-80 mesh size sample (0.18 to 0.25 mm). This was done
to minimize non-uniformities in the velocity-profile (and, hence, the contact
time) of a gas flowing through the packed bed which can arise due to two
mechanisms. The first is simply the diffusion of the species in the axial direc-
tion of the flow. For most gases, the ratio of diffusional to convective (or bulk
flow) transport increases as the Reynolds number of the fluid in the packed
bed decreases. For a given gas flowing above a Reynolds number of 1 [6] this
ratio is less than 1. The present system, 1,000 ppm methanol in nitrogen flowing
at 200 cm3/min through a packed bed mean particle diameter of 0.21 mm, has a
Reynolds number greater than 2 and is, therefore, in the convective transport
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3. EXPERIMENTAL
dominated regime, thereby limiting non-uniformity in contact time due to dif-
fusion.
The second mechanism of dispersion in a packed bed reactor is due to a non-
uniform bulk velocity profile. In particular, it has been shown that in tubular
reactors of small diameter, Dt, with respect to the particle size diameter, D , the
voidage, or fraction of cross sectional area which is devoid of solid particles,
increases dramatically near the wall of the reactor. This effect is due to the in-
terference of the wall with the packing of the solid material. A general criterion
states that the ratio of tube diameter to particle diameter should be above 20.
In the present case, we have a tube diameter of 4 mm with an average particle
size of roughly 0.2 mm resulting a Dt/D = 20.
Hence, the physical conditions of the experimental microreactor were chosen
to minimize the dispersion effects of both axial diffusion and non-uniformities
in the bulk velocity profile.
An additional concern of exothermic reactions studies is whether the heat
produced by the reaction is sufficient to raise the reactor temperature
significantly. In such cases the furnace temperature can read one value while
the local catalyst temperature (which is the temperature of the gas at the
reaction site) can be somewhat higher. This would lead to inaccurate
conclusions concerning the required temperature for a given extent of conver-
sion of the methanol. Since the microreactor was operated in an electrical
furnace, it could be assumed that the reactor operated under nearly adiabatic
conditions. An estimate of the adiabatic temperature rise across a packed bed
reactor is given in Appendix B. Using the values for our system, the tempera-
ture rise across the reactor can be calculated to be roughly 53° C for 30 percent
methanol conversion when a pure methanol feedstock is used. This would be
unacceptably high since that temperature rise could induce further reactions of
the DME to subsequent hydrocarbons as described in reaction (2). The same
calculation with even complete conversion of a 1,000 ppm methanol feedstream
in nitrogen shows an adiabatic temperature rise of only 0.18° C. Hence, it was
determined to use a dilute (1,000 ppm) mixture of methanol and nitrogen for
the screening test.
Experimental Procedure
Each catalyst was crushed and sieved. The 60-80 mesh cut was-calcined.in a
muffle furnace at 300° C for 1-1/2 hr. A 0.40 cm I.D. pyrex glass tube approxi-
mately 62 cm long was used for the reactor vessel. A small amount of glass
wool was compressed and forced into the reactor to form the support for the
catalyst bed. Sufficient catalyst was then added to form a 4 cm long catalyst
bed after gently tapping the sides to induce particle settling. The final mass
varied from catalyst to catalyst due to differences in particle densities. Another
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3. EXPERIMENTAL
compressed plug of glass wool was added to contain the catalyst particles. The
glass tube containing the 4 cm catalytic bed was then placed in the reactor fur-
nace and connected to the incoming and outgoing lines. A complete schematic
of the experimental apparatus is given in Figure 2.
For the reasons described above, a 1,000 ppm methanol in nitrogen stream was
used as the feedstream for this phase of the work. Sufficient back pressure was
made available for the system by setting the bottle regulator at 80 psi during
each run. The methanol was metered from the bottle through a Tylan FC-260
mass flow valve. The valve was controlled by a Sierra "Flow Box" mass flow
controller. The methanol then passed directly to a 10 port gas chromatographic
sampling valve (see Figure 3). The sampling valve provided an in situ
calibrating procedure for the methanol feed stream as follows, (with reference
to Figure 3):
First, the methanol coming from the vaporizer (which was not needed in Phase
1 since a gaseous feed stream was used) entered the 10-port-sampling valve.
There it was routed through a bypass loop, leaving the valve to be fed to the re-
actor inlet (see Figure 2). After passing through the catalyst bed, the stream re-
entered the 10-port-sampling valve as labeled in Figure 3 from the reactor.
When the sampling valve was positioned as in the configuration to the left in
Figure 3, the stream passed directly to the vent and was routed to the laborato-
ry flue. The sampling valve then switched to the configuration shown on the
right in Figure 3. At this time, the helium carrier gas for the gas chromatograph
swept the sample out of the bypass loop to the gas chromatograph for analysis.
At the same time, the reacted gas stream from the reactor entered the 10-port-
sampling valve to be routed through a reactor loop of the same size as the
bypass loop and then to vent. While still in this configuration, the unreacted
methanol stream from the vaporizer was routed directly to the reactor inlet
without passing through the bypass loop. After the bypass gas sample analysis
by the gas chromatograph was completed, the valve was switched back to the
configuration on the left in Figure 3. Now the helium carrier gas swept the
sample out of the reactor loop passing it on to the gas chromatograph for analy-
sis. In this manner, both inlet and outlet streams could be alternately
monitored, thereby correcting for daily fluctuations in GC performance.
A Varian Model 3700 Gas Chromatograph equipped with a flame ionization
detector (FID) was used for hydrocarbon analysis. Separation was effected by
a 1% SP -1,000 60/80 carbon packed column. The injector temperature,
column temperature and detector temperature were maintained at 80°, 60°,
and 200° C, respectively. Peak area integration was calculated by a Hewlett-
Packard 3392A Integrator. Since each injection of the gas sample to the gas
chromatograph required 5 minutes for the complete elution of'all the reaction
components, the valve was switched every 5 minutes. This was effected by a
Hewlett-Packard 19405A Sampler-Event Control Module. Using a prepro-
RTI 3-5
-------
o
o
o
o
c
re
-C
Mass Flow VA A
Control Valve /Y^*
Thermocouple
Readout
I I
Selector
O
Furnace
Temperature Controller
Liquid Pump
(Phased)
Mass Flow
Control Box
Vaporizer
Temperature Controller
y
y
Vaporizer
(Phase II)
Vent
1i
10-Port
Sampling Valve
Reactor Furnace
rooooooooooooooi
Gas
Chromatograph
u
m
X
•o
m
CO
en
Figure 2. Microreactor catalyst testing facility.
m
I
-------
To
Reactor Inlet
To Gas
Chromatography
Vent
JO
H
Figure 3. 10-Port sampling valve arrangement.
He
From
Vaporizer
To Reactor
Reactor Inlet
To Gas
Chromatography
Vent
m
x
•o
m
3J
m
•z.
-------
3. EXPERIMENTAL
grammed method, the 3392A integrator controlled the 19405A sampler-event
module. The ten-port sampling valve was housed in a Carle Model 4301 Valve
Oven configured with a Carle Model 4200 Valve Actuator.
Since the experimental testing required analysis at a variety of reactor
temperatures for a given LHSV, a temperature control unit was needed to
drive the Thermcraft reactor furnace. To this end, an Omega microprocessor
based controller (Model No. CN2011) unit was interfaced with a Tandy 1200
HD personal computer. Software developed by RTI enabled various setpoints
to be stored and fed to the Omega temperature controller at timed intervals.
For a given set point, the Omega temperature controller brought the reactor
furnace to the set point within a ±5° C tolerance. A typical run involved storing
a series of nine set points in the computer and beginning the run. Typically the
set points were 50°, 85°, 100°,115°, 130°, 150°, 175°, 200° and 225° C. The
50° C run provided a pure methanol output to the gas chromatograph which
could be used later for calibration purposes since no reaction occurred.at this
low temperature. Each set point was maintained for 1 hour during which
approximately five samples were taken. Of these, the last three were used for
methanol conversion calculations.
After each run, the peak areas calculated by the 3392A integrator were entered
in a macro-driven LOTUS spreadsheet with the 50° C input and output
methanol peak areas being entered as calibration peaks. The ratio of outlet to
inlet areas (in this case, the 0% conversion case) provides a correction factor
taking into account the pressure drop across the reactor bed. This is required
since the FID is a mass detector rather than a concentration sensitive detector.
Hence, the sample stored in the bypass and the reactor loops (see Figure 3) are
gases of slightly different densities. Although the reactor and bypass loop
volumes are nominally of the same size, different methanol or product masses
are injected in each case to the GC. Even when no conversion is occurring
across the catalyst bed, the total amount of methanol passing through the flame
ionization detector will be less in the case of the reactor loop since the gas there
-is of lower density due to the pressure drop through the catalyst bed. By
entering the input and output peak at 0% conversion (50° C) that ratio can be
used to correct the data for the rest of the run since this density difference
should remain essentially constant. This was done automatically by the
LOTUS spreadsheet.
Once the data were entered, two plots were produced: one showing the metha-
nol conversion as a function of reactor temperature, the other showing approxi-
mate concentrations of both methanol and dimethyl ether in the product
stream. These plots are discussed in Section 4.
RTI 3-8
-------
4. RESULTS
SECTION 4
RESULTS
The results of each catalyst test are presented in a graph depicting the conver-
sion of methanol as a function of reactor temperature. Figures 4 through 10
show these conversion plots for the alumina catalysts tested. Similarly, Figures
11 through 17 present the results for the zeolite catalysts while Figures 18
through 25 present results for the metal oxide catalysts. The equilibrium con-
version of methanol to dimethyl ether was shown in the RTI July 1987 proposal
to be approximately 90% under these conditions. However, small amounts of
DME will react further to higher products, thereby allowing more methanol to
be converted.
For selected catalyst tests, a mixture of 290 ppm dimethyl ether in nitrogen was
fed to the reactor at room temperature to provide DME calibration areas.
Subsequent results could then be represented in terms of actual ppm quantities
of methanol and DME in the product stream. These data are shown in Figures
26 through 30 for selected catalysts. An approximate material balance is
possible in that,
2CH3OH
(6)
produces one part of DME for every two parts of methanol consumed. The
data in Figures 26 through 30 are presented in terms of parts per million by
volume, so multiplying the DME levels by two and adding them to the metha-
nol level should yield roughly the 1,000 ppm level of methanol feed. At higher
levels of conversion, it is likely that some conversion of the DME to higher hy-
drocarbon occurs which is not detectable at these levels using present equip-
ment. It is significant to note that at a methanol conversion of roughly 30% the
material balance is quite accurate, indicating that virtually all the methanol
converted at these lower levels forms DME only.
This result can be used to develop a meajure to rank the catalysts since a
product stream containing 15% by volume of DME in unreacted methanol is
required to start an engine under cold start conditions and a 15 volume %
DME stream corresponds to 30% conversion of methanol. As the data in
Figures 25-30 would indicate, substantially all of the. methanol converted at
these lower levels of conversion forms DME rather than the higher hydrocar-
bons with a concomitant formation of coke. Hence, selectivity at these
minimum conversion levels is not an important criterion for catalyst selection.
The data presented in Figures 4-25 were then used to identify the reactor tem-
perature required to convert 30% of the feed methanol for each catalyst. That
temperature, T30, was then used to rank the activity of the candidate catalysts.
These results are given in Table 9. .
RTI 4-1
-------
4. RESULTS
I
o
HI
0
c
0
o
i
o
II
5
c
0
o
100 120 140 160 180 200
220
240
Temperature
Figure 4. MeOH conversion, AL-5207.
80 100 120 140 160 180 200 220 . 240
Temperature
Figure 5. MeOH conversion, AL-5307 E 1/16.
Rtl 4-2
-------
4. RESULTS
I
o
I
"5
0
o
0.1 -
80
100
120
1 i r
160 180
i 1 1 r
200 220
240
Temperature
Figure 6. MeOH conversion, AL-5407.
x
o
a
'o
c
o
o
0.9 -
0.8 -
0.7
0.6
0.5
0.4
0.3
0.2
0.1
-O |
80
100
i i i
120
180
200
140 160
Temperature
Figure 7. MeOH conversion, AL-3945.
220
240
RTI 4-3
-------
4. RESULTS
1
i
o
01
c
0
o
0.9 -
0.8 -
0.7
0.6
0.5
0.4
0.2
0.2 -
0.1 -
o I o—i 9 r~s~T v 1 1 1 1 1 r 1 1 1 1 1—
80 100 120 140 160 180 . 200 220 240
Temperature
Figure 8. MeOH conversion, AL-3996 R.
1
o
0)
2
'o
c
o
O
0.9 -
0.8 -
0.7 -
0.6 -
0.5 -
0.4 -
0.3 -
0.2 -
0.1 -
0
-Q—
i t i i i i i i i T i n n i i i
80 100 120 140 160 180 200 220 240
Temperature
Figure 9. MeOH conversion, CS 331-1.
RTI 4-4
-------
4. RESULTS
i
o
a
o
o
o
I
o
c
0 .
o
.80 100 120' 140 160 180 200 220 240
1
0.9 -
0.8 -
0.7
0.6
0.5
0.4
0.3
0.2 -
0.1 -
80
Temperature
Figure 10. MeOH conversion, CS331-4.
100
My i
120
i i i
140
160
180
200
220
240
Temperature
Figure 11. MeOH conversion, LZ-20.
RTI 4-5
-------
4. RESULTS
i
o
a
c
0
o
X
o
0)
2
c
0
o
80
100
200
220
240
Temperature
Figure 12. MeOH conversion, LZ-Y-72.
100
120
140
160
180
200
220
240
Temperature
Figure 13. MeOH conversion, LZ-M 8.
RTI 4-6
-------
4. RESULTS
I
o
01
2
'o
c
0
o
1
0.9 -
0.8 -
0.7 -
0.6 -
0.5 -
0.4 -
0.3 -
0.2 -
0.1 -
i i i i i i i i t i i i i i
80 100 120 140 160 180 200 220
240
Temperature
Figure 14. MeOH conversion, Z6-06-02.
1
r
o
c
0
o
0.9 -
0.8 -
0.7 -
0.6 -
0.5 -
0.4 -
0.3 -
0.2 -
0.1 -
0
100
120
200
220
140 160 180
Temperature
Figure 15. MeOH conversion, Z6-06-02,B6980.
240
RTI 4-7
-------
4. RESULTS
i
o
2
'o
c
0
o
180
200
220
240
Temperature
Figure 16. MeOH conversion, MCG-7.
i
o
2
'B
c
0
o
80
100
120
180
200
140 160
Temperature
Figure 17. MeOH conversion, MCG-8.
240
RTI 4-8
-------
4. RESULTS
X
o
2
'o
c
0
u
0.1 -
80
100
120
180
~\ 1 1 1 r
200 220 240
Temperature
Figure 18. MeOH conversion, ZR-0304T.
I
o
c
o
0
0.9 -
0.8
0.7
0.6
0.5
0.4
0.3
0.2 -
0.1 -
0
1 1
80 100
—i 1 1 1 1 1 1 1 1 1 1 1 1—
120 140 160 180 200 220 240
Temperature
Figure 19. MeOH conversion, TI-0720.
RTI 4-9
-------
4. RESULTS
i
o
c
0
o
I
o
o
C
0
o
1
0.9 -
0.8 -
0.7
0.6
0.5
0.4
0.3
0.2 -
0.1 -
0 -I-B
80
100
120
140
160
180
200
220
240
Temperature
Figure 20. MeOH conversion, T-312.
200
Temperature
Figure 21. MeOH conversion, T-314.
220
240
RTI 4-10
-------
4. RESULTS
i
o
9
2
c
0
o
140 160 180 . 200 220 240
80
• • • Temperature
Figure 22. MeOH conversion, T-317.
i
i
o
ai
2
c
0
o
0.9 -
0.8 -
0.7
0.6
0.5
0.4
0.3
0.2 -
0.1 -
-B-
80
100
120
180
200
HO" 160
Temperature
Figure 23. MeOH conversion, T-1502A.
220
240
RTI 4-11
-------
4. RESULTS
I
o
5
0
c
0
u
1 -
0.9 -
0.8 -
0.7 -
0.6 -
0.5 -
0.4 -
0.3 -
0.2 -
0.1 -
0 -
8
— o-— * "a-
-_O D '
— n — i n — rr^^
n I 9 \^J I I I I I I I ' • I I 1 "1 1 ' • •
0 100 120 140 160 180 200 220 240
X
o
c
o
o
Temperature
Figure 24. MeOH conversion, T-1502B.
0.1 -
80
100
120
140
Temperature
Figure 25. MeOH conversion, 7913-S K-306.
RTI 4-12
-------
4. RESULTS
1.1 -
1 -
0.9 -
0.8 -
a 0.7 -
u"""
3= 0.6-
Q D
(A
o J °-5 ~
? 0.4-
5
0.3 -
0.2 -
0.1 -
0 -
e
~~~^^8^lv
\
\
\
V ^ +_ +_
x^*
0 100 120 140 160 180 200 220 240
Temperature
Q MeOH + DME
1.1 -
.1 -
0.9 -
0.8 -
a 0.7 -
I! ae-
(/i
o J 0.5 -
? 0.4-
2
0.3 -
0.2 -
0.1 -
0 -
8
Figure 26. MeOH & DME outputs, A L -5207.
^°^\
\
1
\
\
A
/ \
^^^ -a — a
i i i i i i i i i i i i i i i i
3 1 00 1 20 140 1 60 1 80 200 220 240
Temperature
D MeOH + DME
Figure 27. MeOH & DME outputs, AL-5407 E 1/16.
RTI 4-13
-------
4. RESULTS
1.1
£
a
a
o
"
o
01
2
1 -
0.9 -
0.8 -
0.7 -
0.6 -
0.5 -
0.4 -
0.3 -
0.2 -
0.1 -
0
1—i 1 1 1 1 1 1 1 1 1 r 1 1 1 1 1—
80 100 120 HO 160 180 200 220 240
Temperature
D MeOH •*• DME
Figure 28. MeOH & DME outputs, LZ-20.
£
a
a^
J c
a o
§|
it
o
01
2
i . t
.1 -
0.9 -
0.8 -
0.7 -
0.6 -
0.5 -
0.4 -
0.3 -
0.2 -
0.1 -
. 0 -
LJ~X\
\
\
\
a
%
\
\
\
\
\
\ i
V "•" "~^"™™"™""~— | - L .
y*a _ ~_+
/ \ ~~~
/ \
S/ \ g^ a—
~*
i i i i i i i i i i i 1 1 i 1 1
80 100 120 140 160 180 200 220 240
Temperature
D MeOH + DME
Figure 29. MeOH & DME outputs, M-8.
RTI 4-14
-------
4. RESULTS
160 180 200 220
80
100 120
140
Temperature
D MeOH + DME
240
Figure 30. MeOH & DME outputs, Z6-06-02 1/16.
RTI 4-15
-------
4. RESULTS
Table 9. Ranking of All Candidate Catalysts Tested According to T3Q
Catalyst Catalyst family T3Q (°C)
LZ-M8
CS-331-4
AL-5407
MCG-8
CS 331-1
AL-5307
AL-3945
AL-5207
LZ-20
MCG-7
Z6-06-02
AL-3996R
LZ-Y-72
7913-S K306
T-317
T-314
ZR-0304T
Z6-06-02, B6980
T-312
Zeolite
Alumina
Alumina
Zeolite
Alumina
Alumina
Alumina
Alumina
Zeolite
Zeolite
Zeolite
Alumina
Zeolite
Metal oxide
Metal oxide
Metal oxide
Metal oxide
Metal oxide
Metal oxide
118
131
135.5
136.5
138
142
142.5
142.5
145
145
146
148.5
164
171.5
178
186
188
215
216.5
Catalysts Which did not Convert 30% of the Feed Methanol
Under the Conditions of the Screening Test
TT-0720 Metal oxide
T-1502B Metal oxide
T-1502A Metal oxide
RTI 4-16
-------
5. DISCUSSION
SECTION 5
DISCUSSION
The results given in Table 9, in which all the catalysts tested are ranked
according to their activity for 30% conversion of the feed methanol, are useful
in separating some of the candidate catalyst families. In particular, the metal
oxide catalysts are less active than either the alumina or zeolite. Even those
catalysts in which the metal oxide is supported by an alumina or zeolite
(e.g., Z6-O6-O2-B6980) shows significantly higher T30 values than their non-
metallic counterparts (e.g. Z6-06-02).
It is also interesting to note that the T-317 catalyst showed virtually no metha-
nol (see Figure 20) at higher conversions while all of the other catalysts showed
a maximum of roughly 90% methanol conversion. This non-equilibrium value
of methanol conversion on the T-317 catalyst indicates that further reaction of
DME was occurring. Indeed, upon inspection of the gas chromatographic data
some lighter compounds were detected although no identification was made.
Among the aluminas and zeolites no such clear definition exists. The most
active catalyst was found to be the ammonia exchanged zeolite M-8 provided
by the Union Carbide Corporation. It is known that a mixture of acid and base
sites, perhaps provided by the zeolite and ammonia respectively, is necessary
for a favorable dehydration catalyst. The alumina catalysts, CS-331-1 and -4
provided by United Catalyst Incorporated are primarily alumina oxide with
higher surface area than the Harshaw AL-series (see Tables 3-8). Although the
inspection properties of the AL-5207, -5307 and -5407 catalyst are very similar
(see Table 3), other literature not reported here shows that the pore volume
distribution of the three catalysts are slightly different. In particular, the pore
size distribution of the AL-5407 is slightly narrower with the majority of pore
volume attributable to pores less than 150 A Both AL-5207 and AL-5307
exhibit pores larger than 1,000 A Whether this narrow pore size distribution
provides a larger number of adjacent sites in the sense of the unimolecular
mechanism proposed by Swabb and Gates [5] is unknown. The catalyst litera-
ture provided by Harshaw (makers of 5207,5307,5407) does state that this group
of alumina catalysts is designed for dehydration reactions. They go on to say
that these aluminas have a slightly alkaline nature. Indeed all three of these
catalysts provide significant activity.
The catalysts provided by the Mobil Research & Development Corporation
(MCG-7 and -8) were sent without literature so that their composition is
unknown. It is thought that these are examples of the ZSM-5 catalyst used by
Mobil in their MTG second stage reactor wherein the DME is converted to
RTI 5-1
-------
5. DISCUSSION
higher hydrocarbons for gasoline production. No difference between MCG-7
and -8 was indicated by Mobil.
Again, the metal oxide catalysts show a significantly lower activity as evidenced
by T30 values of more than 20° C above the least active alumina. Of particular
note are the T-1502A and B catalysts which are silica supported rather than alu-
mina supported catalysts. Although silica-alumina is highly acidic, silica by it-
self exhibits no acidity. The titanium oxide catalyst 0720 also has no acidic
support and exhibits virtually no activity.
RTI 5-2
-------
6. PHASE II PLAN OF WORK
SECTION 6
PHASE I! PLAN OF WORK
In light of the data presented as Phase I of this contract it is likely that a
suitable commercial candidate exists for the initiation of Phase II. The goal of
this phase is the design and construction of a prototype dehydration reactor ca-
pable of providing a combustible mixture of methanol and DME (> 15% vol.
DME) at a mass flowrate sufficient to start a spark-ignition engine.
Calculations (see Appendix C) show that if an energy flow of 6 kcal/sec is
needed to start an engine, a volumetric flowrate of liquid methanol of roughly
89.8 cm3/min must be fed to the dehydration reactor. This value can be used in
a simple reactor design. A practical constraint on the design of the prototype is
one of size. The need to mount this unit in the engine compartment of most
vehicles limits total reactor volume. To illustrate this point a maximum unit
size will be assumed as a tube 10 cm I.D. and 35 cm in length. If the catalyst
were to occupy 70% of the internal volume of this reactor, its LHSV is
3
T rrqy _ 5388 cm (liquid methanol) /y>
hr (0.70 (ir) (5 cm)2 (35)
= 2.8 hr"1
This should be compared with the LHSV of the microreactor system used in
Phase I:
200 cm3
min
60 min
hr
1000 cm MeOH3
10
6
3
cm
total
22
mol
,400
cm (liql
3
cm
0.
025
mol
LHSV =
(0.70) (ir) (.2 cmr (4 cm)
= 6.08 x 10"2 hr (8)
It is seen that a LtlSV of more than two orders of magnitude greater than that
used in Phase I will be needed in the prototype. For 30% conversion, the feed
methanol will have to be at a higher reactor temperature. Referring to the data
published by Chang et al. [1] shown in Figure 1, the space time (1/LHSV)
required for the prototype (1/2.8 hr"1 = 0.36 hr) is greater than the space time
needed for 30% conversion (roughly 10"3 hr) measured by those investigators.
This means that the temperature, T30, for the prototype unit must be higher
than the 150° C measured in the Phase I but lower than the 371° C studied by
Chang et al.
RTI 6-1
-------
6. PHASE II PLAN OF WORK
Some problems exist in simply raising the reactor temperature for 30% conver-
sion at the higher LHSV in the prototype. The methanol dehydration reaction
is considerably exothermic. Therefore, the heat that evolves during the conver-
sion of methanol in the reactor must be carried away at a sufficient rate to
prevent a significant rise in the reactor temperature. Otherwise, the higher
temperature will raise the reaction rate and, consequently, the heat of reaction,
which thereby liberated, leads to yet higher temperatures. The maximum
temperature, Tad, is that which would result in an adiabatic system and can be
calculated using a standard expression as given by Kramers and Westerterp [7].
The calculation (given in Appendix D) shows that Tad would equal 377° C.
Thus, if the prototype reactor is allowed to operate adiabatically, the gas phase
temperature will rise approximately 177° C over the length of the unit. Exactly
where the rise will occur is not readily calculated although some approximate
results will be discussed shortly.
The selectivity of the methanol dehydration will suffer dramatically if such a
temperature rise occurs. It was shown in equation (7) that the LHSV required
to start a cold engine is 2.8 hr"1 or a space time (1/LHSV) of 0.36 hr. This is
much longer than the approximately 10"3 hr Chang et al. used to achieve a 30%
conversion of methanol at 371° C on ZSM-5. Hence, a significantly lower reac-
tor temperature would suffice in the case of the automobile prototype (say,
200° C). If, on the other hand, such a unit experienced a 177° C temperature
rise at its midpoint, the remainder of the reactor would operate at 377° C. The
space time for this reactor would be 0.36 * 2 = 0.18 hr. From Figure 1, a re-
actor operated at 371° C and space time of 0.18 hr will produce primarily
higher hydrocarbons. This, in turn, will lead to significant coking with
subsequent catalyst deactivation.
Of course, the prototype will not be operated adiabatically. The temperature
rise of the gas phase reactants will then depend on a balance between the rate
of heat production (via reaction) and the rate of heat transferred away from the
reactor. Probably the simplest configuration of such a non-adiabatic methanol
dehydration reactor is shown below:
*• Gaseous Products
Cold Methanol Fuel
Vaporization
This reactor utilizes the methanol fuel as a coolant while vaporizing and
preheating it as reactant to the dehydrator. The heat requirements for this
process will be described below. For now, however, it is important to focus on
the reactor itself. Gaseous methanol will enter at a temperature of roughly
-------
6. PHASE II PLAN OF WORK
200° C. The dehydration reaction will be initiated thereby generating heat. It
was shown above that a temperature rise of 177° C is possible with no heat
exchange resulting in the rapid deactivation of the catalyst. As the rate of heat
exchange is increased (from zero), however, the maximum temperature, Tmax,
reached in the reactor will drop. At sufficiently high values of heat transfer, the
reaction will proceed in a controlled fashion resulting in prolonged catalyst life
and proper product selectivity. It has been shown [8] that the transition from
controlled operation (with small Tmax) to a rapid and severe temperature rise
(with large Tmax) is a critical phenomenon. Bilous and Amundson [9]
demonstrated the profound effect small changes in the coolant stream temper-
ature had on the gas phase temperature in an exothermic reaction system (see
Figure 31). The region of parameter sensitivity can be estimated without resort
to an exact solution of the relevant heat and mass balances by a method due to
Barkelew [10]. This analysis showed that a relationship between two
dimensionless parameters exists which defines the region of reactor sensitivity.
The first of these parameters is defined as,
N j =
ad
RT2
I c J
' AH '
r1
p
which is related to the adiabatic temperature rise (see Appendix B). Here Tc is
the coolant temperature. The second parameter describes the cooling (heat
transfer) capacity of the system.
(10)
where 1^ is the rate constant as defined in Appendix D, p is the gas density, C
is the specific heat capacity, U is the overall heat transfer coefficient and Rt is
the radius of the tubular reactor.' The relationship between these two
dimensionless parameters is shown in Figure 32.
Using the kinetic data calculated in Appendix D, the values of these parameters
may be computed for the prototype system. The following values will be used:
E = 2.7 x 104 cal/mol
AHv = -2.2 xlO3 cal/mol
Cp = 12.4 cal/mol K
R = 1.987 cal/mol K
Tc = 375 K
U = 5.4 cal/cm2 hr.K
1^ = 1.2 x 104 cm3/cm3 hr (at 140° C)
p = 9.92 x ID'4 g/cm3
Rt = 5 cm
RTI 6-3
-------
6. PHASE II PLAN OF WORK
300
10 20
Residence time
30
Figure 31. Effect of Tc on the temperature profile of a PFTR [9].
RTI 6-4
-------
6. PHASE II PLAN OF WORK
u
(0
2.5-
2.0-
1.5-
1.0-
0.5
0
Insensitive
Sensitive
10 20
50
100
200
N
ad
Figure 32. Relationship between two dimensionless parameters.
Note the assumed coolant temperature, TC, of 102° C. This temperature will
vary in the reactor configuration shown in Figure 5, from the boiling point of
methanol, 65° C, to the reactor inlet temperature. This latter value was
assumed to be 140° C. Tc was assumed to be the average of these two or
102° C. Using these values,
Nad =17
Nc =0.47
= 2.8xlO-2.
This value is clearly in the sensitive region in Figure 32 indicating that a
runaway temperature excursion is possible.
The reactor system described by the above values must be changed.
Unfortunately, not all of the parameters that make up Nad and NC an
ble. In particular, U/Rt and k^, could be effectively varied.
The overall heat transfer coefficient, U, describes the heat transfer from the
-reacting gas through the reactor vessel to the coolant medium (in this case,
liquid methanol). The rate limiting step in this exchange is that of the gas to re-
actor vessel. While little can be done to improve this rate, using the reactor
configuration shown in Figure 31, the surface area over which the heat is trans-
ferred can-be increased. That is, Rt can be decreased. In fact, even if ten 3.2
cm I.D. reactors were used in a shell and tube arrangement (i.e. ten tubes
surrounded by the liquid (boiling) methanol all housed in a larger diameter
vessel), Nc would only be increased to 1.5. The ratio of Nc/Nad would only be
9 x 10~2 which is still clearly in the sensitive region.
RTI 6-5
-------
6. PHASE II PLAN OF WORK
Alternatively, a less active catalyst could be used thereby reducing k^, and
decreasing Nad. Equation (D.7) shows, however, that a decrease in ^ would re-
quire an increase in Vcat for a constant level of conversion and feedrate thereby
requiring a larger reactor vessel. This would be unacceptable from a volume
standpoint.
Apparently a more sophisticated reactor design would be required to improve
U in NC in an effort to de-sensitize the dehydration reactor. A number of
options exist, one of which would be a series of tubular reactors with interstage
cooling.
Catalyst
Feed
V Cold
Methanol
Here, each reactor stage allows the gas phase temperature to rise due to the
heat of reaction. Before the temperature reaches a point which would degrade
product selectivity and catalyst life, however, the reactants pass through a non-
catalyst cooling section before entering a second reactor stage. The tempera-
ture history of the gas-phase reactants would be
si
Js
£ |
»E
o §
Inlet
Condition
Product
Stream
Distance or Time
As the concent-ration of methanol decreases along space (or time), the space
(or time) needed to reach the maximum allowable temperature will increase.
Thus, proper design would provide increasingly longer reactors.
RTI 6-6
-------
6. PHASE II PLAN OF WORK
It should be noted that the optimization of any reactor design will require some
mathematical modeling of the reaction network and reactor system. This will
require some limited kinetic data which can be obtained using the present
microreactor apparatus already operational. By operating in a low conversion
("differential mode"), initial rate data can be measured relatively easily. This
data can then be used in conjunction with the appropriate reactor model to
evaluate its optimal design.
Dissociation Reaction
In the case of the exothermic dehydration reaction, it was shown that the system
could develop a runaway temperature excursion detrimental to both product
selectivity and catalyst life. To maintain a (parameter insensitive) system,
however, a completely novel reactor concept may be sought. One such
concept, originally brought forth by the U.S. Environmental Protection Agency
(EPA), uses an endothermic catalytic reaction which could provide an in situ
heat sink to limit a temperature excursion. The methanol dissociation reaction,
CH3OH + CO + 2H2 (11)
is a highly endothermic reaction yielding CO and H2. These products are suit-
able for cold starting combustion requirements in that they are volatile and
have high octane values. Work previously by EPA dealing with this reaction for
a different application has identified a dissociation catalyst with favorable
properties. Preliminary work with this catalyst was performed using the same
equipment and procedure discussed in Section 3. A 4 cm bed of this dissocia-
tion catalyst was found to have a T30 of 152.5° C (see Figure 33). A 50/50 (by
volume) mixture of the dissociation catalyst with AL-3996R (T30 = 148.5° C)
exhibited a T30 of 150° C (see Figure 34). Although the product gas analyzer
could not detect CO and H2, the two catalysts appear to be compatible in that
the effective activity of the 50/50 mixture as measured by T30 lies between the
T30's of the two pure catalyst beds. Hence, no inhibition was measured.
Such a mixture makes a multi-reaction system possible
2CH3OH + CHgOCHg + HgO •» higher hydrocarbons (12)
CO, H2
Note that conversion of DME to CO and H2 is also indicated. Tests analogous
to the methanol dehydration but using DME demonstrated that the conversion
of DME (presumably via dissociation) was equally feasible (Figure 35).
RTI 6-7
-------
6. PHASE II PLAN OF WORK
i
o
01
2
c
0
o
X
o
1)
c
o
u
I I <-" I I I
160 180 200
80
0.9 -
0.8 -
0.7
0.6
0.5
0.4
0.3
0.2
0.1 -
Temperature
Figure 33. MeOH conversion, dissociation catalyst.
-O-T — "?-
1 1 1 1 r
80 100 120 140 160 180 200 220 240
Temperature
Figure 34. MeOH conversion, dissociation and AL 3996R.
RTI 6-8
-------
6. PHASE II PLAN OF WORK
0.9 -
0.8 -
0.7 -
r
o
I
'o
c
0
0
80
200
220
160 180
Temperature
Figure 35. MeOH conversion, dissociation in DME.
240
The advantage of providing both dehydration and dissociation reaction
pathways lies in relative heating requirements. The dissociation reaction
requires approximately 20 kcal/mol of methanol consumed while methanol de-
hydration releases 2.2 kcal/mol. If a dehydration catalyst .considerably more
active than the dissociation catalyst was mixed with the dissociation catalyst, an
in-situ safety net would be provided. As the dehydration reaction rate
increased due to the liberation of heat, the dissociation reaction rate would be-
gin to utilize more heat in the formation of CO and Hr
A situation would exist which is analogous to the so-called predator (wolf) -
prey (rabbit) system. In this scenario, the rabbit population is kept in check by
an ever-present wolf population. In times of large rabbit proliferation (rapid
dehydration rate) the wolf (dissociation rate) population grows due to the
growing food supply. In turn, the wolf population is held in balance because
the rabbit supply is consumed at a greater rate.
The selection of the appropriate dehydration catalyst (the wolfs food supply)
will depend upon the relative effect of flow rate, heat transfer and both dehy-
dration and dissociation reaction kinetics. To design and construct the
RTI 6-9
-------
6. PHASE II PLAN OF WORK
optimum reactor utilizing a dehydration/dissociation reaction system would re-
quire both the measurement of dehydration and dissociation kinetics and a re-
actor analysis based on the predator-prey model. Many of the latter theoretical
descriptions exist in. the applied mathematics literature so that no new model-
ing development is required. These models provide a steady state analysis
which would be extremely useful in determining relative catalyst mixtures.
Process Energetics
Regardless of whether a pure DME catalyst or a mixture of methanol dehydra-
tion and dissociation catalyst are considered, an ultimate application of this
contract is to design a prototype reactor system capable of starting a cold
engine. If the liquid methanol fuel is initially at a very low temperature (say
-10° C), the cold start apparatus must vaporize this liquid methanol and
preheat it to the initial inlet condition needed for the dehydration reaction
(approximately 200° C). Once in the reactor, the parameter sensitivity
described in the previous section can be important due to the exothermicity of
the dehydration reaction. It may be beneficial in this regard to add some disso-
ciation catalyst. In any event, the heat given off by complete conversion of
methanol to DME is given by:
= 89-8 CI°3 0-7914 g mol 2.2 x 103 kcal
dehydration min cm 32g mol
= 4.9 x 103 .
mm
The amount of heat required to bring the initially cold methanol to the inlet
condition of the reactor (200° C) is given by
Total heat {*£} = h*at ?f ded, to ™ jse + vaporization
liquid methanol to methanol to
(14)
heat needed to , f «
raise gaseous = 25 x 10 ~r~
reactor condition
Obviously, considerably more warm-up heat is consumed than is released by
the dehydration reaction, even at a 100% conversion. Consequently, some ad-
ditional heat must be made available before this apparatus can be used effi-
ciently. This energy could conceivably be provided by the battery of the vehicle
although given this energy availability it could be argued that the dehydration
reactor itself is superfluous. That is, if the battery can provide 20 x 103 cal/min,
it could certainly provide the additional 4.9 x 103 cal/min, thereby delivering a
RTI 6-10
-------
6. PHASE II PLAN OF WORK
stream of vaporized methanol at 200° C to the engine. Even so, this is unlikely
to be feasible; an alternative process should be considered. This process is
shown schematically as:
Warm Gaseous
Products
T
Cold Methanol
Fuel
:> Phase
•>-To Engine
In this system the reactor products are assumed to be vapors at 200° C
consisting of unconverted methanol, water, DME and perhaps a mixture of CO
and hydrogen. The reactor exchanges heat with the incoming methanol feed
stream, preheating it. The products are then passed through a condenser which
liberates a significant amount of heat forming now a two-phase mixture. The
vapor phase of this stream would consist of CO, H2, some DME, water and
methanol. The liquid stream would consist of liquid methanol, some water and
DME. The proportions of DME in the vapor and liquid phases are dictated by
the vapor-liquid phase equilibrium. A simple vapor-liquid equilibrium
calculation is provided in Appendix E for the two component system of DME
and methanol. The results show a liquid stream between 10 - 30° C should
contain a DME mole fraction of roughly 10-20% DME. Since this mixture
would exist under its own vapor pressure in the product phase separator, the
liquid stream could be then pumped to the fuel system (i.e., fuel injectors).
Once the mixture was sprayed in the cylinders or in the air intake, the DME
would readily vaporize producing a combustible mixture suitable for starting
the engine. This is very similar to the concept of winterized gasoline in which
the lighter component of the gasoline is increased during the winter months.
RTI 6-11
-------
7. REFERENCES
SECTION 7
REFERENCES
1. Chang, C.D., and AJ. Silvestri, /. Catal., 47:249-259,1977.
2. Persona] communication between RTI and Mobil Research and
Development Co., 1988
3. Panzar, J., "Characteristics of Primed Methanol Fuels for Passenger Cars,"
SAE paper 831687.
4. Winterbottom, J.M., "Hydration and Dehydration by Heterogeneous
Catalysts," in Catalysis-Specialist Periodical Report, Vol. 4, Royal Society of
Chemistry, London, 141-174,1981.
5. Swabb, E.A., and B.C. Gates, Ind. Eng. Chem. Fund., 11(4), 1972.
6. Levenspiel, O. "Chemical Reaction Engineering," 2nd ed., Wiley, 1972.
7. Kramers, H., and K.R. Westerterp, Elements of Chemical Reactor Design
and Operation, Academic Press, 1963.
8. Froment, G.F., and K.B. Bischoff, Chemical Reactor Analysis and Design,
Wiley, 1979.
9. Bilous, O., and N.R. Amundson, AIChE Journal, 2,117, (1956).
10. Barkelew, C.H., Chem. Eng. Prog. Symp. Series No. 25,55, p. 37 (1959).
RTI 7-1
-------
8. SUMMARY AND FUTURE WORK
SECTION 8
SUMMARY AND FUTURE WORK
The research done in Phase I of this contract shows that there are commercial
catalysts potentially suitable for the methanol dehydration reaction. This is
evidenced by the wide range of dehydration activities as measured by T30 given
in Table 2. Some problems exist, however, in the development of a prototype
dehydration reactor. These were discussed in the Phase 2 plan of work.
Primarily, the exothermicity of the methanol dehydration reaction can result in
a runaway temperature excursion within the reactor. This can lead to loss of
product selectivity by the formation of higher hydrocarbons with the
concomitant formation of coke (carbonaceous material) which can lead to
catalyst deactivation. Although more sophisticated reactor designs can be used
to minimize the likelihood of these excursions, the reaction system is quite
sensitive to slight changes in operating conditions. It is proposed to study the
kinetic rates of both the methanol dehydration reaction over a limited group of
candidate catalysts as well as the methanol dissociation reaction over the
dissociation catalyst explored in Phase I. With this information, a relatively
simple theoretical analysis is feasible by which a dehydration/dissociation mix-
ture reactor could be designed. The dissociation reaction pathway would be
provided only as a heat sink to prevent reactor runaway. Finally, it was noted
that the process energetics would require the condensation of the product
stream in order to allow operation without a significant amount of external
power requirements. The condensed product stream is believed to be suitable
for cold starting an internal combustion engine in a mechanism completely
analogous to winterized gasoline.
It is recommended that the design parameters of a methanol dehydration/
dissociation reactor be pursued to enable the design of a stable cold start
apparatus. To that end the following work is proposed:
• study the reaction kinetics of the methanol dehydration over a select
number of candidate catalysts
• study the reaction kinetics of the dissociation reaction over the dissociation
catalyst explored in Phase I
• complete the reactor analysis required for optimal reactor designs and
catalyst mixture proportions
RTI 8-1
-------
8. SUMMARY AND FUTURE WORK
• complete a more rigorous thermodynamic calculation of the multi-phase
equilibrium formed by water, DME and methanol
It should be noted that the measurement of the chemical kinetic rates is a
relatively simple task for which the equipment used for Phase 1 can be used
with only slight modification (to include CO and hydrogen detection).
Although it is perhaps possible to design a prototype based exclusively on the
methanol dehydration reaction using the data made available through Phase I,
it is felt that the level of sophistication needed to control that reaction would be
high making the expense of the ultimate reactor prohibitively large. Alter-
nately, the dehydration/dissociation concept appears to be inherently stable,
permitting a simpler design.
RTI 8-2
-------
APPENDIX A
APPENDIX A
SOLICITATION FOR COMMERCIAL METHANOL DEHYDRATION
CATALYSTS
Dear
Research Triangle Institute is involved with the developement of a system
for the dehydration of methanol vapor to a mixture of dimethanol ether in
water. The purpose of this research is to overcome cold start problems in
methanol powered vehicles. The ultimate application of the catalyst and
system which we will develop will be for such vehicles. Our immediate goal is
twofold:
• Screening of a number of catalysts that are suitable for the
dehydration of methanol to dimethanol ether.
• The assembly of a system using the best catalyst from our
screening test to be tested on a stationary methanol powered
automobile.
Our review of the literature on this reaction indicates that the
dehydration of methanol is an intermediate step over acid catalyst with the
ultimate products being light olefins and hydrocarbons. However, in contrast
to processes such as the methanol to gasoline process, we wish to develop a
catalyst that will produce dimethanol ether selectively. The literature also
seemed to indicate that desirable catalyst characteristics are strong lewis
acidity, little or no bromstead acidity, and weak lewis basicity. Conditions
of interest for our study are atmospheric pressure, temperatures below about
350*C, and high space velocities.
We would appreciate your recommendations for catalyst that you may produce
which could meet the objectives of our program. Specifically, we would
appreciate your providing us with roughly 100 grams of candidate catalyst that
you feel would work. In return, we will keep you advised of the results of
our screening tests.
We plan to begin our screening in mid-December and would appreciate your
providing samples for our use by then, if possible. We greatly appreciate
your cooperation in this effort.
Please contact me at (919) 541-7272 to discuss any details of our goals
and how your catalyst might be evaluated.
Sincerely,
James J. Spivey, Ph.D.
RTI A-1
-------
APPENDIX B
APPENDIX B
CALCULATION OF ADIABATIC TEMPERATURE RISE IN A
TUBULAR REACTOR
The maximum temperature reached in a tubular reactor occurs when the
reactants are completely converted under adiabatic conditions. This
temperature known as the adiabatic temperature (T ,) can be calculated using
a standard expression as given by Kramers and Westerterp [1].
T £
ad dT = AH /f d£ (B.I)
/
T
where C is the specific heat capacity of the gas stream, TQ is the temperature of
the incoming stream, x is the molar fraction of the reacting species, AHr is the
heat of reaction and f is the fractional conversion. The maximum Tad occurs
when £f = 1 (100% conversion). Although the equilibrium conversion of
methanol to DME alone is limited to 85%, the DME can react further for
higher hydrocarbons. This will allow complete depletion of the methanol.
Equation (B.I) can now be evaluated to yield
AH
Tad - To = -« of *, (B-2)
For pure methanol forming DME (the heat of reaction due to the further
reaction of DME is also exothermic but will be neglected),
Cf = 1
x = 1
AHr = -2.2 x 103
C = 12.4
p mo
T = 200° C (assumed inlet temperature)
Tad is calculated to be 377° C. If a feedstream of 1,000 ppm of methanol in
nitrogen is converted,
RTI B-1
-------
APPENDIX B
Reference
x = 1 x ID'3
which yields a Tad - T0 of 0.18° C.
1. Kramers, H., and K.R. Westerterp, Elements of Chemical Reactor Design
and Operation, Academic Press, 1963.
RTI B-2
-------
APPENDIX C
APPENDIX C
CALCULATION OF METHANOL REQUIRED TO START AN ENGINE
For the preliminary design of a prototype catalytic dehydrator, an estimate of
the required methanol throughput is needed. To this end, the combustion
energy required to start a spark ignition engine was assumed to be 6 kcal/sec
(EPA). To satisfy cold starting requirements a product stream of 15% DME is
needed [1] which corresponds to 30% conversion of methanol. To provide 6
kcal/sec assuming 15 mole % DME, the volumetric flow rate of gas from the
dehydrator can be found as follows:
o
Q _ 6 kcal/sec 22.400 cm
* ~ E x. (-AH .) mol
3. Cl
where
Q = volumetric flowrate of gas from the dehydrator, cm3/sec
x- = mole fraction of component i, dimensionless
AH • = heat of combustion of component i, calculated herein at 25 ° C with
1 all reactants and products as gases, kcal/mol
For our system
AHC(CH3OH) =-160.3 kcal/mol
AHC (CH3OCH3) = -331 kcal/mol
AHc(H,O) =0 kcal/mol
Substituting
Q = 829 cm3 STP/sec
To result in 829 cm3 STP/sec of gas containing 15% DME, 15% FLO, and 70%
methanol, the volume of liquid methanol, Qm, that must be fed to the
dehydrator is thus:
3 3
n - goo cm mol gas cm (liquid) 32 g 60 min
m ~ sec 22,400 cm3 STP °'7914 g ™* min
3
Q = 89.8 ££_ liquid methanol
xm mm ^
Reference
1. Panzar, J., "Characteristics of Primed Methanol Fuels for Passenger Cars,"
SAE paper 831687.
RTI C-1
-------
APPENDIX D
APPENDIX D
ESTIMATION OF GLOBAL KINETICS OF METHANOL DEHYDRATION
It is possible to use data presented in Section 3 to estimate the overall chemical
kinetics of the methanol dehydration reaction,
2CH3OH * CH3OCH3 + HgO (D.I)
at an approximation for use in preliminary design. At low conversions (< 50%)
reaction D.I is far from equilibrium (see Section 2) so that only the forward
reaction should be considered. For heterogeneous reaction systems a first
order expression based on the volume of catalyst (m cat) can be assumed to
approximate this reaction's rate,
r = kyC (D.2)
where r is the rate of methanol conversion (m3/hr m3cat), k^, is the rate constant
(m3/hr m3cat) and C is the concentration of methanol (M) in the feed (fd)
(m3M/m3fd).
In estimating the rate constant from data measured in the tubular microreactor,
an analysis similar to that given by Kramers and Westerterp [1, p. 25] will be
used. That analysis shows that, at steady state, the extent of conversion, £, of a
reactant varies with distance along the axis of the reactor, z, according to,
(D.3)
where Scat is the cross-sectional area of catalyst and $v is the volumetric
flowrate of methanol (m3 hr) fed to the reactor at reactor conditions. This
expression can be rearranged and integrated to give
^ idf . /"'S^d, (D.4)
o o
where f, is the extent of conversion at the reactor outlet and V „ is the total
31 Cat
volume of catalyst in the reactor. According to reaction (D.I) the gas mixture's
density should remain constant. The rate expression (D.2) can be rewritten in
terms of the extent of reaction as.
RTI D-1
-------
APPENDIX D
r = kyCo (1-0 (D.5)
where C0 is the concentration of methanol in the feedstream at the reactor
inlet. Equation E.I can now be integrated to yield,
v o
This expression can be solved for k^,,
- ef) (°-7)
o cat
For the catalyst screening tests described in Section 3 the following values apply
0V = 1.2 x ID'5 rr^/hr
Co =1.0xlO-3m^/m3fd
Vcat = (1 - e)Vr = (1 - 0.3)(5.03 x 10'7 m3) = 3.52 x 10-7m3cat .
Here Vr is the total reactor volume (4 mm I.D., 40 mm long) and e is the
approximate void fraction in the reactor.
In this manner the first order rate constant can be calculated for any catalyst
tested under Phase I. For the purpose of the preliminary design work
presented here, the results of catalyst AL-5207 (Harshaw) presented in
Figure 28 will be used:
(1.2 x 10~5(m2/hr)
k = -- 3 o 3 - - - n-t - ln(l-0.25) (D.8)
v (1.0 x 10 <5(mu/mdfd)(3.52 x 10 m cat)
M
The rate constant is a function of reactor temperature
ky = Ae~E/RT (D.9)
where A is the pre-exponential factor, E is the activation energy of the reaction
on the particular catalyst, R is the gas constant and T the temperature in
Kelvin. The activation energy can be estimated by plotting the In k vs 1/T for a
number of reactor temperatures, as follows:
RTI D-2
-------
APPENDIX D
JL
o.i
0.2
0.25
0.3
3.6 x 103
7.6 x 103
9.8 x 103
1.2 x 104
Ink
8.2
8.9
9.2
9.4
1/T(K4)
2.51 x lO'3
2.46 xlO'3
2.44 x ID'3
2.42 x lO'3
10.0 - -
9.0--
8.0--
2.4 2.42 2.44 2.46 2.48 2.5
1 (K~1) X 103
The slope of this curve is, then, E/R = 13.6 x 103 K. The apparent activation
energy E, of this reaction on AL-5207 is approximately 27 kcal/mol.
Reference
1. Kramers, H., and K.R. Westerterp, Elements of Chemical Reactor Design
and Operation, Academic Press, 1963.
RTI D-3
-------
APPENDIX E
APPENDIX E
CALCULATION OF THE VAPOR-LIQUID EQUILIBRIUM (VLE)
DATA FOR A TWO-COMPONENT MIXTURE OF DME AND METHANOL*
A first approximation to the vapor-liquid equilibrium relationship for a binary
mixture of dimethyl-ether (DME) in methanol (MeOH) at 1 atm pressure is
given by:
For DME, YDM£ = KDM£ XDM£ (E.I)
where y is the vapor-phase mole fraction, K is the proportionality constant to be
determined, and X is the liquid phase mole fraction.
For MeOH, YMeQH = KMeQH XMcQH (E.2)
The addition of (E.I) and (E.2) yields
XDME
Two methods will be used:
Method 1: Assume Henry's Law for
Assume Raoult's law for
KDME = 7°°DME PDME
MeOH
The infinite dilution activity coefficient for DME, 7°°DME, was calculated using
UNIFAC (Universal Quasi Chemical Functional-Group Activity Coefficient
model) while Psat, the vapor pressure of both components, was calculated by
Antoine's equation.
The VLE results are given in Table E-l for both 10° C and 30° C.
* Source: Tony Rogers, RTI.
RTI E-1
-------
APPENDIX E
Table E-1
Temperature (°C)
10
30
7«DM
1.47
1.5
psat
DME
3.81
7.08
XDME(%)
16.8
7.6
Method 2: Assume Raoult's law for both DME and MeOH
y _ pSa,t
rMeOH
The VLE results are given in Table E-2 for both 10° C and 30° C.
Table E-2
Temperature (°C)
10
30
psat
DME
3.81
7.08
XDME(%>
24.7
11.5
The actual liquid-phase DME concentration will likely be between the
extremes shown in the above tables. Because the target reactor conversion of
methanol is roughly 30%,othe DME concentration at the phase separator will
be approximately 15'%'.° T^ais js;substantially above infinite dilution of DME in
methanol, so the Raoult's law-results are thought to be more accurate. Note,
however, that the proportionality constant, K, is concentration dependent and a
more involved iterative procedure should be used to predict XDME with greater
accuracy.
RTI E-2
------- |