Report on the Construction and Operation of a Mars Methanol in situ
Propellant Production Unit
Robert Zubrin, Tomoko Kito, Brian Frankie
Pioneer Astronautics
445 Union Blvd., Suite 125
Lakewood, CO 80228
303-980-0890
Introduction
This report describes work accomplished on the Mars methanol in situ propellant
production (MMISPP) project, contract number NAS 9-97082. This project involved
design and construction of a demonstration oxygen/methanol production facility sized to
fuel the Mars Sample Return (MSR) mission. Project work was carried out from March
through September 1997. John Connolly was the JSC program manager and Robert
Zubrin was the principal investigator at Pioneer Astronautics.
During the MMISPP project, Pioneer successfully built and operated two chemical
synthesis units representing the cores of machines capable of manufacturing oxidizer and
a variety of fuels out of primarily indigenous Martian material. The fuel and oxidizer can
be used for internal combustion or rocket engines.
During the MSR mission, the
spacecraft can use the technology demonstrated during this project to fuel the Earth
return vehicle while the rover collects samples on the Martian surface. The MMISPP
project has some resemblance to the Martian in situ resource utilization project carried
out at Martin Marietta in 1993 – 1996. The previous project methanated carbon dioxide
via the Sabatier reaction using an analogous process configuration. The system also
electrolyzed water produced during the reaction.
This system, called the
Sabatier/Electrolysis (S/E) system can achieve a mass leverage of 10.3 and also can be
applied to the MSR mission.
The units designed and built during the MMISPP project include a Reverse Water Gas
Shift (RWGS) unit and a Methanol Synthesis (MEOH) unit. A process flow diagram of
the units built in the Phase I project are shown inside the dotted box in Figure 1. Outside
Page 1 of 46
of the dotted box in Figure I is the proposed scope of work for the Phase II project. The
RWGS unit works in the following manner: Liquid hydrogen is transported from Earth
to Mars, where it is combined with carbon dioxide acquired from the Martian atmosphere
in a catalytic reactor to produce carbon monoxide and water. Water is condensed and
separated from the gas phase. Unreacted feed components are compressed, recovered
from the gas phase in a membrane unit separator, and recycled to the catalytic reactor.
Effluent from the membrane unit (“retentate”) is sent to the MEOH unit.
At H2/CO2 mixture ratios of 1:1 nearly all the hydrogen is reacted to make water, which
can then be electrolyzed to produce oxygen and hydrogen, which can be recycled. Used
in this way, the hydrogen brought to Mars can be recycled many times to produce an
enormous amount of oxygen and CO. Alternatively, the feedstock to the RWGS can be
run with an excess of hydrogen, in which case the effluent from the RWGS will contain
both CO and H2. Such a mixture is known as synthesis gas and is the ideal feedstock for
making methanol, dimethyl ether, or higher hydrocarbons.
Page 2 of 46
Figure 1: Process Flow Diagram of end to end MMISPP system. The portion inside the
dotted box was completed in the Phase I project. The portion outside of the dotted box is
the proposed Phase II work.
The MEOH unit works in the following manner: Effluent from the RWGS unit,
consisting primarily of carbon monoxide with some residual unreacted hydrogen, is
combined with fresh hydrogen feed, if required, and sent to a catalytic methanol synthesis
reactor.
The synthesis reactor combines one carbon monoxide molecule with two
hydrogen molecules to produce methanol.
Production from the methanol synthesis
reactor is condensed and separated from the gas phase. Unreacted feed components are
separated in a second membrane separator and recycled to the RWGS reactor feed. The
second membrane retentate, consisting primarily of excess carbon monoxide, is vented
from the system.
The MEOH unit was designed in a generic fashion so that different varieties of catalytic
reactors could be tested without changing the process configuration. In addition to the
Page 3 of 46
methanol synthesis reactor, Pioneer experimented with a hybrid methanol/DME reactor
and a Fischer-Tropsch hydrocarbon synthesis reactor.
Picture of the completed RWGS and fuel synthesis units
It should be noted that the general process configuration and system design of the MEOH
unit has been an established industrial technology for more than 80 years, and thus is
considered open art. However, Pioneer’s work on this project was, to our knowledge, the
first specifically tailored to drive the RWGS to near 100% completion, thereby making it
a viable candidate technology for in-situ propellant production in the Martian
environment.
Page 4 of 46
Accomplishments during the Phase I MMISPP project
Pioneer recorded an impressive list of accomplishments during the MMISPP project.
Highlights include:
1) Development, manufacture, and demonstration of a catalyst which is 100% selective
for the RWGS reaction at a wide range of conditions.
2) Design, construction and operation of a MMISPP machine including a RWGS unit
and a methanol synthesis unit.
3) Operation of the RWGS machine in oxygen production mode and attained mass
leverages in excess of 250.
4) Discovery that by altering the reactor temperature, pressure, and feed ratio, the
RWGS unit could be run in combined Sabatier/RWGS mode with potential mass
leverage of 20. Achieved an actual mass leverage during operation of 16.5, which
compares to a 10.3 leverage for the S/E unit.
5) Operation of the machine in a mode to produce a combined 50/50 molar CH4/CO
ratio fuel with a stoichiometric oxygen ratio. In this mode, the system demonstrated a
mass leverage of 31 with a 23 excess oxygen mass leverage.
6) Demonstration of production of synthesis gas (syngas) feed for methanol, dimethyl
ether (DME, =CH3OCH3), Sabatier, or Fischer-Tropsch reactors. The quality of
syngas produced was sufficient to allow a methanol/O2 leverage of 16.3 or a FischerTropsch/O2 leverage of 22.4.
7) Demonstration of production of a 79% methanol/ 21% water fuel product with no
other detectable contaminants.
8) Demonstration of conversion of 8% of the feed carbon dioxide to dimethyl ether
(DME) in a one pass (no recycle) hybrid reactor.
9) Demonstration of conversion of at least 44% of the feed carbon dioxide the ethane or
higher hydrocarbon species in a one pass Fischer-Tropsch reactor.
Page 5 of 46
10) Demonstration of complete recovery of gaseous hydrogen in a two membrane loop
system, with no gaseous hydrogen detected in the system vent.
Development and Production of RWGS catalysts
Literature Survey on RWGS Catalysts
A literature search on Chemical Abstract was performed in order to find catalyst
candidates for the RWGS reactor. The focus of this literature search was to determine
which catalysts are the most selective towards production of carbon monoxide. For this
application selectivity of the catalyst is more important than its activity. Based on this
search, three groups of catalysts appear to be suitable for this application:
1. Cu supported catalysts
2. Au supported catalysts
3. Mo compounds
The Cu supported catalysts have shown good activity and outstanding selectivity to
produce CO from CO2. For example, Nozaki et al. (1987) reported that their Cu/alumina
catalyst demonstrated 28% CO2 conversion with 100% CO selectivity when the reactor
was operated at 350 °C under atmospheric pressure with SV (space velocity) of 100
ml/min/g-cat and CO2/H2 feed ratio of 1/4. The loading of Cu was 12 wt%. Even though
they have tested various metal supported catalysts on alumina under the same conditions
(Ni, Rh, Ru, Pt, Pd and Re), the Cu catalyst was the only one that exhibited exclusive
selectivity to CO. The other metal catalysts tended to produce more methane or, in some
cases, only methane.
5 wt% Cu/silica catalyst was also able to give at least 97% CO selectivity (Kitayama,
1997). The catalyst was evaluated at 350 °C with a feed ratio CO2/H2 of 1/4. The
conversion to CO was 60% under a pressure of 150 torr. The catalyst activity was
Page 6 of 46
improved by adding a little amount of Ni to Cu while maintaining high activity.
However, if the catalyst had too much Ni, it started forming more methane. The authors
recommend Ni0.1Cu0.9/silica as the best catalyst in this series.
When 5 wt% Cu/silica catalyst was operated at 60 bar and 280 C with a feed mixture
(CO2: 22.7%, H2: 67.2%, Ar: 10.1%) and SV of 50 ml/min/g-cat , the selectivity to
carbon monoxide was decreased and more methanol was produced (Dubois, 1992).
According to their results, the catalyst showed 17% CO2 conversion and carbon
monoxide and methanol selectivities of 76 and 24%, respectively.
The selection of support material for Cu catalysts and the reaction pressure seem to be
two critical parameters affecting selectivity. For methanol synthesis process Cu/ZnO
catalyst is usually employed under 10 bar with the temperature range of 250 - 350 C.
The Cu/ZnO catalyst is more selective to produce methanol than carbon monoxide even
at atmospheric pressure (Fujita, 1992), which implies that ZnO plays an important role in
the reaction chemistry.
Au supported catalysts on metal oxides were tested by Sakurai (1993). They used a
hydrogen-carbon dioxide feed mixture with argon (CO2: 23.4%, H2: 66.2%, Ar: 10.4%)
under 8 atm with a space velocity of 3000 ml/h/g-cat. The temperature range evaluated
was 150 - 400 C. Among those they tested, Au/TiO2 and Au/Fe2O3 were found to have
good selectivity and a conversion level close to the thermodynamic equilibrium value. At
400 C 35% carbon monoxide and 3.3% methane were produced on Au/TiO2, and 38%
and 1.3% on Au/Fe2O3.
However, the Au/Fe2O3 catalyst tended to produce more
methanol in the temperature range of 150 - 300 C than Au/TiO2.
According to the recent publication by the same group (Sakurai 1997), by decreasing
reaction pressure from 50 to 1 bar, CO selectivity of the Au supported catalysts were
significantly improved. For example, CO, methanol and methane were produced with
selectivities of 86, 4 and 10%, respectively, on 2 atom% Au/TiO2 at 50 bar while more
than 99% of CO was formed at 1 bar on the same catalyst. The most remarkable property
Page 7 of 46
of the Au catalysts is that they are able to reach CO2 conversions that are close to the
equilibrium limit even at temperatures as low as 250 K. The major disadvantage of gold
catalysts would be cost compared with copper and other materials.
Mo catalysts have attracted some attention for the RWGS reaction. Saito and Anderson
(1981) tested bulk Mo compounds for CO2 reduction and they found that Mo metal had
fairly higher activity than MoS2. On the other hand, MoS2 supported on TiO2 appeared to
demonstrate the best performance in the MoS2 supported catalysts (Taoda, 1991). 13%
CO2 conversion was achieved with more than 99% selectivity on this catalyst at 400 C
with a CO2/H2 feed ratio of 1. The conversion of the catalyst at 400 C was much lower
than that at thermodynamic equilibrium. One advantage of using sulfide catalysts is that
the catalyst can't be deactivated by sulfur compounds present in the feed.
If no sulfur exists in the feed stream, use of Mo oxide catalysts is more practical. The
MoO3/ZnO catalyst was tested at 873 K with a CO2/H2 ratio of 1 (Suzuki, 1995). The
CO2 conversion was 30% with close to 100% CO selectivity. The other RWGS catalyst,
NiO/ZnO, showed higher activity (38%) but selectivity dropped to 93%. With excess
amount of CO in the feed stream at 903 K, NiO/ZnO failed to demonstrate good
performance because of carbon deposition and methanation. On the contrary, under the
same reaction conditions, the MoO3/ZnO catalyst maintained close to 100% CO
selectivity. The function of the ZnO support in this study was not explained.
From the above information it appears that Cu supported catalysts on alumina or silica
are the primary candidates for the RWGS reaction. Au supported catalysts are worth
trying if necessary but less desired do to considerations of cost and preparation
procedures.
Mo oxide supported catalysts also have a good chance to achieve the
requirements for the RWGS reactor.
Catalyst Evaluation
Page 8 of 46
Experimental Apparatus
Catalyst evaluation tests were conducted in order to provide kinetic information of the
candidates for reactor design purposes. The reactor employed was a continuous downflow micro catalytic reactor.
High purity hydrogen and carbon dioxide cylinders
equipped with water removal cartridges were used to feed the reactor. The feed mixture
was preheated to 150 °C before flowing over the catalyst bed composed of catalyst and
quartz chips as a diluent. The catalyst bed was mounted in 1/2 inch stainless steel tube
with a small piece of 100 mesh wire cloth to support the catalyst bed. The bed was
located in the isothermal zone of the reactor. A Lindberg furnace heated the reactor. The
temperature was monitored in the middle of the bed by a J type thermocouple inserted
from the top of the reactor. The reaction products flowed through a heated line to the gas
sample valve in an SRI gas chromatograph (GC, thermal conductivity detector). The GC
column was 10 ft by 1/8 stainless steel tubing and packed with Porapak N (80-100 mesh).
Equipment Calibration
To ensure accuracy, all major pieces of equipment were calibrated. In addition, a blank
run was made on the stainless steel tube containing only quartz chips at 400 °C to verify
that it has no activity for the RWGS reaction.
Gas Chromatograph
The compounds expected in the product stream were determined by their relative
retention times. To obtain clear peak separation the initial oven temperature of 50 °C was
ramped at 20 °C per minute to 150 °C where it was held for 10 minutes. The response
factors of the compounds were determined with a gas mixture containing 25% of each,
H2, CO, CH4, CO2. Water and methanol response factors were determined with their
liquid mixtures. Accurate analysis of hydrogen requires spiking the GC helium carrier
gas with hydrogen, therefore, 10% H2 in He was used as a carrier gas for GC. Water
eluted in a very broad peak with poor reproducibility. The water concentration was
Page 9 of 46
therefore determined from the concentrations of the other products by using reaction
stioichiometry.
Catalyst Selection
Based on the literature survey, the Cu supported catalysts on g-alumina and silica were
chosen. The loading of Cu was approximately 10 wt%.
Dr. Tomoko Kito brewing up a batch of catalyst
Catalyst Preparation
Supports used were g-alumina (Norton, 1/16’ spheres, BET surface area = 200 m2/g) and
silica gel (Davison Chemical, Grade 57, crashed into 20/40 mesh, BET surface area =
300 m2/g). The support materials were calcined at 500 °C overnight to drive off all the
adsorbed water. After cooling, the support was impregnated with a solution of cupric
nitrate by incipient wetness technique. The impregnated material was then dried at 110
Page 10 of 46
C overnight and calcined at 500 C for 2 hours. The pore volume of each support was
determined prior to the impregnation, and was fund to be 0.60 ml/g for g-alumina and
1.16 ml/g for silica. To ensure filling of the pores, the impregnation was done under
vacuum.
Test Procedure
The catalyst bed consisted of 0.5 g of catalyst and 1.5 g of quartz chips. The catalyst was
reduced in situ using hydrogen at 400 °C for 2 hours, then, the inlet gas was switched to a
hydrogen and carbon dioxide gas mixture at selected flow rates. The H2 to CO2 ratio was
1 and reaction temperature was fixed at 400 °C under atmospheric pressure.
The
operating conditions were maintained constant until a minimum of three consecutive
samples of product stream were reproducible (steady state).
Test Results of Cu/g-Alumina Catalyst
The Cu/g-alumina catalyst was evaluated at the total feed flow rates of 10.4, 19.6, and
42.0 ml/min in this sequence. The results are listed in Table 1.
Table 1. Test Results of Cu/g-Alumina Catalyst
Total feed flow rate
Conversion of CO2 %
Selectivity of CO %
10.4
19.6
100
19.6
13.0
100
42.0
7.97
100
H2 : CO2 = 1 : 1
[ml/min]
Page 11 of 46
As expected, the lower feed flow rate resulted in a conversion approaching the
thermodynamic equilibrium value (24% at 400 °C). It is important to emphasize that no
by-products, such as methane or methanol, were detected throughout the entire run. The
activity and selectivity of this catalyst were satisfactory for the RWGS catalytic reactor.
Approximately 500 g of this catalyst were manufactured and about 250 g were used to
load the reactor.
Design, Construction, and Operation of the Demonstration System
System Overview
System design, procurement, and construction occurred from April through July 1997.
Design was performed with a simulator developed for this application. Material take offs
from the simulator allowed the procurement and construction to proceed quickly and with
little wasted effort. A simplified engineering schematic of the as built system is shown in
Figure 2.
Page 12 of 46
CV1
FV03
CV2
CV3
CV6
CV4
FI3
PLV1
PI1
SP
FV07
FV05
FV11
PI2
FI1
SP
FI2
SP
RWGS
unit
membrane
FV01
FV02
Heat
Exchanger
CO2 H2
Feed Feed
FV10
H2
Feed
FV06
FV12
Compressor
SP
Excess CO
to Vent
MEOH
unit
membrane
CV5
Water
Condenser
PLV2
RWGS
Reactor
SP
SP
LI1
Methanol
Condenser
Methanol
Reactor
TI1
LI2
FV04
Figure 2: Schematic of as-built system
TI3
FV09
The system consists of a fresh feed inlet from bottled hydrogen and carbon dioxide gas.
This fresh feed is mixed with recycled gas from both the RWGS unit membrane and the
MEOH unit membrane, warmed to reactor temperature and passed through the RWGS
reactor catalyst bed. Effluent from the RWGS reactor is cooled, first in the feed/effluent
exchanger and then in the condenser. The condensed water is then phase separated in a
vessel. The aqueous phase can be drawn off the phase separator or simply accumulated
in the vessel.
In the Phase 2 portion of this project, the water will be sent to an
electrolysis unit that will produce pure oxygen for liquefaction and hydrogen for recycle
to the RWGS unit feed. Vapor from the phase separator is compressed to high pressure
and sent to the membrane unit, which recovers hydrogen and carbon dioxide in the low
pressure permeate and rejects carbon monoxide and excess hydrogen in the high pressure
residue that is sent to the methanol converter.
The feed to the methanol converter is made up with fresh hydrogen, if required, and
heated to approximately 220 degrees C. The methanol reactor is filled with commercial
Page 13 of 46
Cu/ZnO on alumina methanol synthesis catalyst. Effluent from the methanol converter is
cooled in the methanol condenser. Condensed methanol product is separated from the
unreacted gas, which then flows to the MEOH unit membrane.
In the membrane,
unreacted hydrogen along with some CO flows through the permeate and returns to the
RWGS unit feed. The membrane retentate, consisting of excess CO, is vented from the
system.
The MEOH unit was designed so that it could flexibly use different catalysts in the
reactor bed to produce a range of potential fuels from the RWGS effluent, including
methanol, methane, Fischer-Tropsch hydrocarbons, and dimethyl ether. Several reactors
were built to examine the production of each of these fuels. As time permitted, the
methanol reactor was removed from the system and the alternate reactors were tested.
Characteristics of Unit Operations
The compressed schedule of the Phase I project meant that off-the-shelf equipment was
used as much as possible. This impacted the operability of the unit in several ways.
Reactor and condenser vessels
The reactor and condenser vessels used 300 cc stainless steel sample cylinders. A pass
through fitting in the top allowed an axial feed tube to penetrate to the bottom of the
vessel. The vapors would then flow up the annular region of the vessel and out the top of
the vessel. A pass through fitting in the bottom of the reactor vessels allowed the
insertion of a Type J thermocouple into the middle of the catalyst bed to monitor reactor
temperatures.
Compressor
The methanol reaction is favored by high pressures, which requires a high pressure
generator in the system. A number of compressors and several configurations were tried
before settling on the process configuration shown in Figure 2. This configuration used a
pneumatically driven positive displacement compressor manufactured by Haskel, Inc.
and driven by the laboratory air supply. The compressor is entirely mechanical; it uses
no electrical components and consumes no electrical power. One of the reasons this
Page 14 of 46
compressor was used was to demonstrate this technology, which could be used on Mars
using the waste gas from the CO2 sorption pumps to provide the driving force. However,
since the pump has a relatively slow cycle rate and a rather large pumped volume,
pressure fluctuations in the system tended to be larger than desired. A 500 cc surge
volume was added immediately downstream of the compressor, which smoothed the
pressure fluctuations to an acceptable level.
Haskel pneumatically driven process compressor
Membranes
Two varieties of membranes were used during the operation of the unit. A Permea Alpha
PPA-22 unit is a standard laboratory air separation unit, and was used for the entire
duration of the RWGS unit operation. This membrane has approximately 5 times the
required surface area for our desired MSR mission production rates. In addition, the
selectivity of the membrane is not as good as one designed specifically for the
application. In operation, the large area allowed the recovery of almost all of the CO2
and most of the hydrogen. Depending on the exact operating conditions, a substantial
Page 15 of 46
amount of CO would also go into the permeate.
However, even with these
disadvantages, the separation in the membrane was good enough that the RWGS unit
routinely achieved very high conversions.
A hydrogen recovery membrane was custom made by MEDAL, Inc. for the MEOH unit
membrane separation.
This membrane has a hydrogen to CO selectivity of 200,
considerably higher than the Permea membrane, which allows it to recover almost all the
unreacted hydrogen in the methanol condenser vapor. However, this membrane was
small enough that it caused a substantial pressure drop in the system. To avoid the
pressure drop, the methanol unit MEDAL membrane was replaced with a second Permea
unit.
Machine Mass
The as built machine has a number of areas where mass could be reduced. The following
table shows the mass of the components in the current machine and the potential for
reducing the components in the next generation machine.
Components
Present Machine Mass
Estimated Next Generation
(kg)
Machine Mass (kg)
RWGS reactor vessel
1
1
MEOH reactor vessel
1
1
RWGS condenser vessel
0.7
0.7
MEOH condenser vessel
0.7
0.7
Compressor
15
5
Compressor surge vessel
1
1
RWGS membrane
0.5
0.3
MEOH membrane
0.5
0.3
Page 16 of 46
Miscellaneous piping,
5
2
25.4
12.0
fittings, and instrumentation
Total
The physical dimensions of the present machine are as follows:
RWGS unit: 1.2 m tall x 0.5 m wide x 0.2 m deep = 0.12 cubic meters
MEOH unit: 1 m tall x 0.4 m wide x 0.2 m deep = 0.08 cubic meters
The unit currently has a large number of extraneous fittings, sample points, valves, etc.,
and additionally has several oversized unit operations, so the actual physical dimensions
of a flight ready system should be expected to be smaller by a factor of 1.5 to 2 in every
dimension.
Unit Operation
History
The RWGS unit was first started on 22 July 97. Operation in integrated RWGS/methanol
mode was first achieved on 11 August 97. Between the start of the RWGS unit and the
end of the program, Pioneer conducted an extensive testing program of the machine.
During a total of 21 running days, over 250 data points were taken. However several
major system reconfigurations were made during August, and the system did not reach its
final configuration until the first week of September. To reduce the data to a manageable
level and maintain a relatively consistent set of baseline conditions, 25 runs, primarily
from the last few days of the testing process, were chosen as most representative of the
different operating conditions and trends found during running the RWGS unit. The 25
RWGS runs are shown in Table 2. For the methanol and fuel synthesis unit, individual
runs are used and individually described as required.
Page 17 of 46
Temperature and flow control unit for RWGS and MEOH units
Page 18 of 46
Table 2: Data from RWGS unit sorted into pressure classes and by temperature
Reactor Conditions
Pressure
Temp
(Bar)
(Deg C)
Feed
Ratio
(H2/CO2) Number
Product compositions (Mole %)
H2
CO2
CO
CH4
H2O
Reactant Conversions
Total
Total Hydrogen
Hydrogen CO2
to Water Oxygen
Leverages
Methane
CO
Total
1.2
1.5
2.2
0.8
0.8
350
350
350
400
450
1.25
1.23
1.38
0.95
1.49
24
23
25
18
9
26.1
18.4
30.4
9.0
5.0
14.8
8.2
14.5
11.4
1.0
29.6
36.7
27.5
39.8
38.0
0.0
0.0
0.0
0.0
6.0
29.6
36.7
27.5
39.8
50.0
0.53
0.67
0.48
0.81
0.93
0.67
0.82
0.66
0.78
0.98
1.00
1.00
1.00
1.00
0.81
9.1
16.0
7.2
35.2
23.5
0.0
0.0
0.0
0.0
2.8
15.9
28.0
12.7
61.6
31.3
25.0
44.0
19.9
96.8
57.6
5.0
4.3
4.6
5.0
5.0
3.6
380
395
400
400
430
445
1.25
1.51
1.04
1.44
2.21
1.29
20
15
19
13
7
8
11.4
25.5
2.1
7.9
2.2
0.0
3.4
5.2
2.3
1.0
1.7
0.5
40.9
34.6
46.6
39.3
23.1
42.6
1.1
0.0
0.8
4.2
16.6
4.8
43.2
34.6
48.2
47.6
56.3
52.2
0.80
0.58
0.96
0.88
0.98
1.00
0.93
0.87
0.95
0.98
0.96
0.99
0.95
1.00
0.97
0.85
0.63
0.84
25.3
10.9
106.3
23.5
12.7
43.6
0.7
0.0
1.7
2.1
3.8
4.0
42.0
19.0
180.0
33.9
9.2
62.3
68.0
29.9
288.0
59.4
25.7
109.9
6.3
6.3
7.0
7.0
6.3
5.7
7.0
330
370
370
370
380
400
432
1.88
1.64
1.12
1.58
2.30
1.49
2.51
4
5
2
3
10
6
1
32.4
20.2
3.1
19.8
35.6
10.8
5.4
3.4
0.6
0.5
1.8
0.7
1.1
0.9
31.1
36.9
46.7
36.5
28.9
38.4
19.8
0.7
1.8
1.0
1.8
2.0
3.8
18.0
32.4
40.5
48.7
40.1
32.9
45.9
55.9
0.51
0.69
0.94
0.69
0.51
0.83
0.94
0.90
0.98
0.99
0.96
0.98
0.98
0.98
0.96
0.92
0.96
0.92
0.89
0.86
0.61
7.7
13.6
76.0
13.7
6.6
20.0
10.8
0.2
0.6
1.6
0.6
0.4
1.6
3.5
12.9
21.7
127.4
21.9
10.2
29.2
6.7
20.7
35.9
205.0
36.3
17.3
50.9
21.0
9.1
10.8
9.1
9.1
9.4
9.8
8.1
335
360
370
380
380
410
425
1.98
1.49
1.98
1.30
1.50
2.37
1.54
12
11
22
17
21
14
16
26.1
14.1
26.9
11.2
19.5
4.4
2.4
1.9
1.7
1.3
2.8
4.3
0.4
1.0
30.4
37.9
31.3
40.4
35.4
22.5
36.8
3.7
2.8
3.1
1.7
1.8
16.7
7.7
37.9
43.5
37.5
43.8
39.0
55.9
52.2
0.63
0.78
0.62
0.81
0.69
0.95
0.97
0.95
0.96
0.96
0.94
0.90
0.99
0.98
0.84
0.89
0.86
0.93
0.91
0.63
0.77
9.0
17.6
9.1
24.0
13.5
11.8
23.6
0.9
1.1
0.8
0.9
0.6
3.5
3.5
12.7
26.8
13.2
38.8
21.4
8.3
29.1
22.6
45.5
23.0
63.7
35.5
23.7
56.2
Out of the 25 runs listed in Table 2, the five runs listed in Table 3 provide the best
examples of operating conditions of different RWGS modes.
Table 3: Runs best demonstrating different RWGS operating modes
Operation
Mode
Oxygen
Run
19
Reactor
Reactor
Feed
RWGS
Pressure
Temperature
Ratio
Net Outlet Leverage Leverage
RWGS
(Bar)
(Deg C)
Chemistry
Leverage
4.6
400
1.04
Production
Methane/O2
7
5.0
430
2.21
Bipropellant
Page 19 of 46
2.1 H2
2.3 CO2
46.6 CO
0.8 CH4
48.2 H2O
2.2 H2
1.7 CO2
23.1 CO
16.6 CH4
56.3 H2O
Fuel
180 CO
Oxygen
Total
106
288
13
26
1.7 CH4
9.2 CO
3.8 CH4
Methane/CO/ 8
3.6
445
1.29
Oxygen
Propellant
Syngas
for 3
7.0
370
1.58
Methanol
production
Syngas for 4
FischerTropsch
Production
6.3
330
1.88
0.0 H2
0.5 CO2
42.6 CO
4.8 CH4
52.2 H2O
19.8 H2
1.8 CO2
36.5 CO
1.8 CH4
40.1 H2O
32.4 H2
3.4 CO2
31.1 CO
0.7 CH4
32.4 H2O
62 CO
44
110
14
36
7.7
21
4.0 CH4
22 CO
0.6 CH4
13 CO
0.2 CH4
Power Usage
During most of the running conditions, power usage for the unit was minimal. RWGS
thermal power required was typically about 60 Watts, which was in fairly good
agreement with the theoretical prediction of 43 Watts from the simulator. Power usage
did increase as recycle ratio increased, with the 60 W figure typical for a 10 mole recycle
to 1 mole fresh feed ratio, which was the preferred operating condition. Electrical power
for the RWGS compressor unit was 0 Watts when the pneumatically driven pump was
used. Earlier in the program, an electrical pump was used, which required about 90
Watts. However, this pump was oversized for the design requirements, and the discharge
usually had to be throttled to reduce the recycle flow rate to a manageable rate.
Estimated required pump power for the desired recycle ratio, if a properly sized
electrically powered pump is used, is about 50 Watts. Thermal power for the methanol
reactor is theoretically 0, since the reaction is exothermic and should supply all the
required heat. However, due to the thickness of insulation used, actual thermal power for
the methanol reactor was typically about 20 Watts. Adding insulation to the reactor
would reduce the methanol power requirement to negligible levels. In any case, the
power requirements for water electrolysis will clearly dominate the total power required
to operate the RWGS and fuel production units.
Production Rate
Page 20 of 46
The production rate of the RWGS unit had no difficulty meeting the desired goals for the
MSR mission of 500 g of bipropellant per 10 hour day. A large number of runs were
performed at 1.5 times of the nominal rates or more, without any noticeable capacity
restraints on the unit. Of course, as capacity goes up, the power usage increases.
General Operation Notes
On the first day of RWGS operation, 22 July 97, the unit achieved 38% conversion of
feed CO2 into carbon monoxide. As experience with the machine increased, these
conversions increased rapidly. By the end of the program conversions of greater than
95% were routine.
Dr. Robert Zubrin operating the RWGS unit
A number of trends are evident during operation of the machine. Figure 3 shows a trend
of increasing carbon dioxide reactant conversion to carbon monoxide with increasing
Page 21 of 46
pressure at 350 and 400 degree C isotherms. In addition, it is clear from this figure that
CO2 conversion also increases as the temperature increases. Both of these results are
expected from theory.
Temperature increases both the equilibrium constant of the
RWGS reaction as well as the reaction kinetics. Pressure does not affect the reaction
equilibrium, but it does increase kinetics by increasing the reactor space velocity. In
addition, a higher system pressure will reduce the vapor fraction of water in the
condenser overhead gas. Finally, higher pressure increases the permeation through the
membrane, meaning more CO2 is forced through and back to the reactor.
Figure 3: CO2 Conversion vs. Pressure
CO2 Conversion (fraction)
1.00
0.90
0.80
0.70
0.60
0.50
0.00
2.00
4.00
6.00
8.00
10.00
12.00
Pressure (Bar)
T = 350 C
T = 400 C
Trend for T = 350 C
Page 22 of 46
Trend for T = 400 C
Figure 4 is a plot of reactant conversion (both CO2 and H2) vs. the hydrogen to carbon
dioxide feed ratio. Runs 6, 10, and 19, which were used to develop this plot, as these
points all have similar pressures of about 5 bar and temperatures of about 400 C. The
plot clearly shows that as more hydrogen is included in the feed, the amount of hydrogen
converted decreases. However, the amount of CO2 conversion remains fairly high over
all these feed ratios, dropping only slightly as hydrogen approaches a 1:1 feed
stoichiometry.
Reactant Conversion
Figure 4: Reactant Conversion vs. Feed Ratio
1
0.9
0.8
0.7
0.6
0.5
1
1.5
2
2.5
Molar Feed Ratio (H2/CO2)
Hydrogen Conversion
Hydrogen Conversion
Carbon Dioxide Conversion
Carbon Dioxide Conversion
Figure 5 is a plot of the total system mass leverage vs. the carbon dioxide to hydrogen
feed ratio. For this plot, all runs with pressures less than 3.5 bar or temperatures less than
350 C were dropped from the data set because these are not optimal conditions for
maximizing the leverage. The data clearly show an increasing trend and a huge leverage
at nearly stoichiometric feed ratios, when the machine is running in oxygen production
mode. Note that the total leverage is the sum of the oxygen and fuel leverages, which
includes carbon monoxide (for a CO/O2 rocket).
Page 23 of 46
Figure 5: Total Leverage vs. Feed Ratio
Total Leverage (kg /kg Hydrogen)
300.00
250.00
200.00
150.00
100.00
50.00
0.00
0.4
0.6
0.8
Feed Ratio (CO2/H2)
Operation of the RWGS Unit
Operation of the RWGS Unit in Oxygen Production Mode
Operation of the machine in oxygen production mode is most clearly demonstrated by
run number 19 from Table 2. In this run, the feed stoichiometry is nearly 1:1 on a molar
basis and the CO2 and hydrogen are nearly entirely converted to carbon monoxide and
water products. The oxygen mass leverage is 106. This operation mode would be useful
for life support for a manned mission or if the oxygen machine is operating in parallel to
Page 24 of 46
1
a S/E machine in order to produce a stoichiometric feed. In addition, if the CO is
recovered, it can be used to fuel a CO/oxygen rocket engine with a total mass leverage of
288. Alternatively, used in this mode the RWGS could simply generate the oxygen
supply for a mission that brought kerosene or some other fuel to Mars. Since rocket
oxidizer to fuel mixture ratios are generally 3:1 or better, such an application would still
offer considerable mission benefit.
Drawing a sample for chromatographic analysis
Operation of the RWGS Unit in combined Sabatier/Oxygen Production Mode
Although the Cu on alumina catalyst is very selective at moderate conditions, it was
found that at higher pressures and temperatures, it also catalyzes the methanation
reaction. This opens the possibility of operating the RWGS unit at conditions that allow
it to produce methane and exactly enough oxygen to provide a stoichiometric fuel burn.
Run number 7 is the best example of the methane/oxygen bipropellant mode for the
RWGS unit. At these conditions – 430 C and 5 bar – the product of the RWGS unit is
Page 25 of 46
16.6 mole percent methane. To burn this stoichiometrically requires 66.4 moles of
oxygen, for a total ideal leverage of 20. This run did not produce quite enough oxygen,
so the actual leverage is only 16.5. However, this is still a large improvement over the
S/E system, which only produces a 10.3 mass leverage.
The conditions under which the Cu on alumina catalyst will produce significant quantities
of methane are shown in Figures 6 and 7. Figure 6 is a function of methane mole percent
in the RWGS unit outlet vs. temperature.
The three lines on the chart show the
exponentially increasing trend of methane production with temperature, which is due to
the increase in the reaction kinetics at higher temperatures. The lowest line on this graph
is the average of all the points in different temperature ranges. Thus the first point would
be the average of the points between 330 and 360, the second between 360 and 390, etc.
This procedure smoothes some of the outlying points and shows a very well behaved
exponentially fitted curve with an R2 value of 0.98. The other two lines are methane
content with increasing temperatures at different isobars. The data show more scatter,
but the trend is still clearly apparent. In addition, the 9 bar isobar shows a higher
trendline than the one at 6 bar, which is expected if methane production is increasing with
pressure.
Page 26 of 46
Figure 6: Methane Production vs. Temperature
Methane in Exhaust (Mole %)
20
18
16
14
12
10
8
6
4
2
0
330
350
370
390
410
430
Temperature (Deg C)
Average Methane Production
Methane Production at 6 Bar
Methane Production at 9 Bar, H2/CO2 = 2
Average Methane Production
Methane Production at 6 Bar
Methane Production at 9 Bar, H2/CO2 = 2
To confirm the trend that methane content does increase with pressure, Figure 7 shows
methane content vs. pressure at two isotherms. Again, the data exhibit some scatter, but
the linear trend is clearly apparent. This trend agrees with theory because the methane
production should increase as hydrogen partial pressure increase in the feed. Also, the
high temperature isotherm has much higher methane content than the lower temperature
isotherm. Using the data from Figures 6 and 7, a curve which gives ideal bipropellant
stoichiometry can be plotted from about 400 degrees C at 10 bar to about 430 degrees C
at 5 bar.
Page 27 of 46
450
Methane in Exhaust
(Mole %)
Figure 7: Methane Production vs. Pressure
20
15
10
5
0
0
2
4
6
8
10
Pressure (Bar)
Methane Production at 400 C
Methane Production at 400 C
Methane Production above 430 C
Methane Production above 430 C
Operation of the RWGS Unit in combined Sabatier/CO/Oxygen Production Mode
One of the problems of operating the RWGS unit in methane/oxygen bipropellant mode
is that the methane and carbon monoxide produced in the RWGS reactor will have to be
separated. A crude separation can be performed fairly simply either with a selective
membrane or with a cryogenic condenser. However, neither of these methods will
produce a high quality methane stream without adding significant complexity to the
process. Therefore, there may be some advantage to operating the RWGS unit to produce
a fuel made of a combination of methane and carbon monoxide at some desired ratio that
can be achieved with a crude separation process.
Run number 8 is one that may be useful for such an operating mode. This run has 4.8
mole percent methane and 42.6 mole percent monoxide. If a 50/50 methane/monoxide
molar ratio fuel is desired, only 90% of the monoxide has to be rejected from the RWGS
reactor outlet stream, which can be accomplished with a single pass through a selective
Page 28 of 46
membrane or a simple methane condensing process. The 50/50 fuel mixture requires 5
oxygen atoms for a stoichiometric burn with a CH4/CO pair, which gives a mass leverage
of 31. In addition, there are 23 excess oxygen mass units left over. If the oxygen can be
used for life support or other purposes, the total mass leverage of this system is 54.
Operation of the RWGS Unit to Produce Syngas
Syngas is commonly used on Earth as the feed to methanol, dimethyl ether (DME), or
Fischer-Tropsch hydrocarbon synthesis reactors. Terrestrial syngas is typically produced
from the steam reforming of coal or natural gas. This process makes it very difficult to
control the hydrogen/carbon ratio in the syngas. For example, a natural gas feedstock
will always be hydrogen rich, while a coal feedstock is typically hydrogen poor. The
RWGS reactor allows flexible adjustment of the syngas stoichiometry for any desired
synthesis reaction, as shown by the data collected during the RWGS unit operation.
Toasting the first batch of Martian Spring water from the RWGS unit
Page 29 of 46
Run number 3 produced a near optimal syngas feed for a methanol or methanol/DME
hybrid reactor. The 19.8 mole percent hydrogen if reacted to completion will yield 9.9
moles of methanol per mole of feed, which requires 29.7 moles of oxygen for a
stoichiometric burn. The total mass leverage of this run producing a methanol fuel is
16.3 and there is an excess of oxygen, which can be used for life support.
Alternatively, run number 4 exhibits an ideal syngas for Fischer-Tropsch hydrocarbon
synthesis. The 32.4 moles of hydrogen will ideally react with 16.2 moles of carbon
monoxide to produce a total mass leverage of 22.4.
Operation of the Fuel Production Unit
Production of methanol
A large methanol peak in the liquid collected in the methanol condenser
Page 30 of 46
Methanol synthesis is favored by very high pressures. The methanol unit was operated at
the highest pressure the machine was capable of generating. System components limited
this pressure to 16 bar, which is near the minimum required for methanol synthesis.
Despite the difficulties generating and maintaining pressure, the unit produced 79% pure
methanol with a water balance. No other impurities were detectable.
The methanol unit was run in full recycle mode with permeate from the methanol
membrane unit returning to the feed for the RWGS reactor. When run in this manner, no
free hydrogen could be detected in the methanol unit vent line, although there was a small
amount of methane. This indicates that all the hydrogen in the unit feed is converted to
useful fuel or to water.
Production of DME
There is currently extensive interest in industry in developing technologies to produce
and distribute DME as a fuel. DME has a very high cetane number and can be used
directly as a clean burning fuel in diesel engines with minor modifications to the fuel
injectors. Production of DME requires the dehydration of two methanol molecules.
Technologies currently being developed use a single catalyst bed to simultaneously
synthesize and dehydrate methanol at less severe conditions than methanol synthesis
requires.
Pioneer constructed a DME synthesis catalyst, made a reactor bed and tested it on 13
September 97. The highest yield achieved in the gas phase sample from the reactor
effluent was approximately 3 mole %. On a carbon basis, this indicates that 8% of the
feed carbon reacted to form DME. The reactor was operating only on a once through
basis (no recycle) at the time, so an 8% yield is quite acceptable.
Production of higher hydrocarbons
The Fischer-Tropsch reaction to create higher hydrocarbons from syngas was discovered
in 1923 and has been the subject of continuing research to the present. Pioneer found a
reference to Fischer-Tropsch activity in a commercial ammonia catalyst and prepared a
Page 31 of 46
reactor using this catalyst. The reactor was tested on 14 September 97, and produced up
to 13% unknown hydrocarbons of C2+ or greater. On a carbon basis, at least 44% of the
feed carbon was converted to higher hydrocarbons in the single pass through the reactor.
After the day’s run, the Fischer-Tropsch condenser was drained and found to have a
clearly visible hydrocarbon layer on top of the aqueous phase.
Conclusion
The MMISPP project has definitively demonstrated that it is possible to build a high
leverage RWGS unit using Martian in situ resources which is capable of operating in
oxygen production mode, methane/oxygen bipropellant production mode and syngas
production mode. The syngas stoichiometry can be adjusted to provide feed for a unit
that can produce methanol, DME, or higher hydrocarbons. Power consumption of the
RWGS was found to be small compared to the irreducible power requirement of the
water electrolyzer needed to electrolyze the product (60 W for RWGS vs 200 W for the
electrolyzer at 50 gm/hr production rate) This means that a RWGS based Mars in-situ
propellant production system promises to be energy efficient. Built out of simple catalyst
beds in steel tubes operating at moderate temperatures, RWGS based systems promise to
be both more robust and scalable than the competing methods of oxygen production
based on ceramic zirconia membranes operating at 1000 C. The efficiency, robustness
and flexibility of the RWGS thus open the way to a number of mission design options
that may dramatically lower the cost and/or increase the performance of the both the
robotic MSR mission, as well as future manned missions. We therefore recommend that
research and development of RWGS based Mars in-situ resource utilization technologies
be pursued aggressively in the future.
Page 32 of 46
Technical Appendix
DME Synthesis
DME (dimethyl ether) is a highly volatile compound with a boiling point of -25 °C.
DME is generally produced by dehydration of methanol over acid catalysts as shown in
equation (1).
2 CH3OH = CH3OCH3 + H2O
(1)
This reaction is much more thermodynamically favored and pressure insensitive than the
methanol synthesis reaction from syngas.
Many solid acid materials have been
investigated as a dehydration catalyst: -alumina, silica-alumina, aluminum phosphatealumina, hydrofluoric acid promoted alumina, and phosphoric acid promoted alumina.
Karpuk and Cowley (1988) reported that hydrofluoric acid promoted alumina catalyst
was highly active for methanol dehydration to ether. The optimized temperature of these
catalysts is known to be between 250 and 350°C. Heteropoly acid catalysts, such as
H3PW12O40 and its salts, are worth trying if a lower operating temperature is required.
The most important advantage of DME synthesis in this application is that it is able to
continuously remove methanol from the system so that the thermodynamic equilibrium
constraint of the methanol synthesis reaction is overcome. As shown in Figure 1, the
methanol synthesis reaction (equation 2) demands very high reaction pressure in order to
have good conversion.
CO + 2 H2 = CH3OH
(2)
At 10 bar the yield of methanol is severely limited by thermodynamics. If methanol
synthesis and dehydration of methanol are combined, the CO conversion at equilibrium
can go up to nearly 40%, compared with 8% in the case of methanol synthesis alone.
Once water is produced by the dehydration reaction, the water gas shift reaction (WGS),
given in equation (3), proceeds rapidly.
Page 33 of 46
CO + H2O = CO2 + H2
(3)
The equilibrium product distribution involving all three reactions is shown in Figure 2.
The overall CO conversion at equilibrium is increased due to the WGS reaction without
having a big impact on the DME production. When the forward WGS reaction occurs at
a very fast rate, the amount of water formed in the syngas reactor becomes negligible.
Therefore, there is practically no requirement for water removal from the fuel condenser.
Solianos and Scurrel (1991) reported their results from syngas conversion experiments to
DME over bifuctional catalysts. According to their study, the combination of methanol
synthesis and dehydration of methanol gave high syngas conversion to form a mixture of
methanol/DME. Their best result was obtained when coprecipitated Cu-Zn-Al catalyst
for methanol formation and -alumina for dehydration was used. The two catalysts were
mixed and placed in one reactor where the highest conversion of CO at 4 MPa, 300 °C
with a GHSV of 16000 hr-1, was 55 - 60%.
Li et al. (1996) evaluated CuO/ZnO/-Al2O3 hybrid catalysts using various preparation
methods. The best results demonstrated 43.7% DME yield (based on carbon mole) with
CO conversion of 63.8% when the reaction was carried out at 270 °C, 3 MPa and 2000
hr-1 GHSV with a stoichiometric feed containing 5% CO2. They concluded that highly
dispersed fine crystallites of CuO/ZnO/-Al2O3 contributed to high DME synthesis
activity.
Fisher-Tropsch (F-T) Synthesis
The reaction in which a mixture of carbon monoxide and hydrogen was converted to
hydrocarbons over iron, nickel, or cobalt catalysts was first discovered by Fischer and
Tropsch in 1923 (Pines, 1981). Those catalysts give different hydrocarbon distributions;
for example, nickel tends to promote methane formation while cobalt promotes carbon
chain growth.
Page 34 of 46
In order to use Fischer-Tropsch synthesis in the Mars in-situ propellant production
project, the fuel produced needs to meet three requirements:
1) have a small H/C ratio.
2) must be stored easily.
3) must not contain heavy hydrocarbons that can cause plugging problems in
the reactor system.
A catalyst for the F-T Synthesis should be selected accordingly. If a F-T catalyst tends to
produce less methane, the first and second issues are mostly addressed. Methane has H/C
ratio of 4, the highest possible number, and its critical temperature is -82.5°C. In terms
of H/C ratio, a more suitable F-T catalyst for this application should form olefins, such as
ethylene (C2H4) and propylene (C3H6), rather than ethane (C2H6) and propane (C3H8).
According to the literature search performed, Fe-based (Snel, 1989; 1988; Huang et al.,
1991; Burkur et al., 1989; 1990) or Fe-Co alloy catalysts (Snel, 1989; Snel, 1988;
Nakamura et al., 1980; Röper et al., 1984) have a good chance to satisfy these factors.
Snel (1988; 1989) had a particular interest in the production of small olefins over
partially degraded iron or iron-cobalt complexes. The hydrocarbon distributions over the
iron complex catalysts are listed in Table 1. In addition estimated H/C values and
pressures required to have complete liquefaction at -50 and -100 °C are given in the same
table.
Table 1 Hydrocarbon Selectivity for Several F-T Catalysts, H/C, and Liquefaction
Pressure (*Data are taken from Snel, 1988; 1989, **Calculated with Aspen Plus)
Page 35 of 46
Selectivity, wt%*
Catalyst
(Temperature = 270°C, pressure C-Fe
C-FeK
C-FeCo
= 2 MPa, Feed H2/CO = 0.5)
Fe
(BASF)
C1 (methane)
8
5
17
19
C2-C5
44
53
55
55
C5-C12
50
43
40
32
C13-C15
8
4
3
3
C19+
2
1
0
1
Olefins in C2
56
76
10
74
Olefins in C3
87
86
72
85
Olefins in C4
84
84
71
81
Olefins in C5
84
83
68
80
Olefins in C6
83
81
66
78
Olefins in C7
81
80
47
72
Estimated H/C**
2.51
2.33
2.89
2.76
at -50 °C
27.4
18.1
42.6
46.8
at -100 °C
8.3
5.4
12.7
13.7
Pressure for liquefaction, bar**
Snel demonstrated that the C-Fe (iron complex) catalyst produced much less methane
(8%) than the BASF Fe catalyst (19%), which is a fused iron catalyst for the F-T
synthesis. The C-FeCo (iron-cobalt complex) showed higher activity , but it tended to
form more methane (17%). One of the least-methane-promoting catalysts was C-FeK
(potassium-promoted iron complex). In the presence of this catalyst only 5 wt% of
methane formation was reported. The promotion with potassium resulted in an increase in
Page 36 of 46
the catalyst activity as well as the selectivity improvement. In order to store this
hydrocarbon mixture, a pressure of 18.1 bar is necessary at -50 °C. An alternative way is
to discard the methane in the mixture and store the rest under a much lower pressure.
Fischer-Tropsch Catalysts
C-73 (United Catalyst, Inc.)
The C-73 catalyst is a fused iron catalyst, similar to the BASF catalyst. According to the
manufacturer’s report, it consists of 67 - 69% Fe, 2 - 3% Al2O3, 0.5 - 0.8% K2O, 0.7 1.2% CaO, and less than 0.4% SiO2. There are a few publications in which this catalyst
is used for the kinetic study of F-T reaction (Zimmerman, et al., 1989; Huff and
Satterfield, 1984).
Modified Recipe for C-Fe
Referring to the C-Fe preparation method by Snel (1988), the recipe for C-Fe was
modified because of the availability of equipment. Ferric nitrate, Fe(NO3)39H2O and
citric acid (HOOC)CH2C(OH)(COOH)CH2(COOH) are dissolved in a minimal amount
of deionized water (approximate weight ratio of ferric nitrate/citric acid = 6.3), and the
solution is placed in a three-neck flask. To concentrate the solution, the flask is kept in a
water bath maintained 60 - 80°C under vacuum using an aspirator unit. It usually takes
24 hours to complete this process. A highly viscous liquid is then moved in a ceramic
bowl, and calcined in a furnace for 1 hour at 400 °C. The catalyst precursor is a light
purple-colored material of low density. A promoter, such as potassium, can be added by
recipient wetness method. The catalyst needs to be reduced in a hydrogen flow at 160 °C
followed by 300 °C.
Page 37 of 46
MTG Process (Methanol-To-Gasoline)
The production of gasoline (a mixture of hydrocarbons, mainly C4 - C10) from methanol
and/or DME over zeolite catalysts was first announced by Mobile Corporation (Meisel et
al., 1976). The catalytic conversion of methanol proceeds at 360 - 415°C and 2.0 MPa as
in equation (1).
x CH3OH = (CH2)x + x H2O
(1)
Methanol is synthesized from CO and H2 (equation 2).
CO + 2 H2 = CH3OH
(2)
Therefore the overall reaction is:
x CO + 2x H2 = (CH2)x + x H2O
(3)
As shown in equation (3), the whole reaction (methanol synthesis/MTG process) is
identical to the F-T synthesis. However, the hydrocarbon distribution formed in the
MTG process is very different from the one in the F-T synthesis. It predominantly
consists of C4 - C10 hydrocarbons including as much as 40 wt% aromatic hydrocarbons
(Chang and Silverstri, 1977). The MTG process tends to make a little amount of light
hydrocarbon gas, such as methane, and heavy hydrocarbons. ZSM-5 class catalysts are
ideal for this process because its aperture diameter of 0.6 nm allows no hydrocarbon
higher than C10 to escape from the zeolite interior (Pines, 1981).
Table 1 represents the effect of space velocity on methanol conversion and hydrocarbon
selectivity as well as fuel H/C ratio and complete liquefaction pressure. The dehydration
of methanol to dimethyl ether is considered to be the first reaction in the MTG process
because DME is a major product with higher flow rate. As LHSV is decreased, the
methanol conversion increases, and when LHSV is 1, the complete conversion of
methanol is reported. The produced hydrocarbon mixture has a H/C ratio of 2.1, even
better than the one with the C-FeK F-T catalyst. In order to store this mixture without
Page 38 of 46
vaporization, 5.6 bar is required at -50°C and 2.0 bar at -100°C, which is exceptionally
low.
There are two options to design the reactor system:
1) Two reactors in series (first reactor for methanol synthesis/dehydration, second
for MTG)
2) Direct synthesis of hydrocarbon from syngas via methanol/DME in single
reactor
In order to choose the second design, development of the catalyst system will be essential
since the optimum catalyst system to do all (methanol synthesis, methanol conversion to
hydrocarbon) is not established yet. However, Cavalcanti et al. (1992) investigated CO
hydrogenation over Pd/zeolite Y to produce methanol, DME, and paraffins. Therefore
metal/zeolite systems may have potential to produce a gasoline-like hydrocarbon mixture
from syngas in single step.
Page 39 of 46
Table Effect on Space Velocity on Methanol Conversion and Hydrocarbon Distribution
at 371 °C (*Data taken from Chang and Silverstri (1977); **Calculated with Aspen Plus)
LHSV
(Volume
of
liquid MeOH/volume
of catalyst - h)
1080 108
1
Product distribution*,
Water
8.9
33.0
56.0
wt%
MeOH
67.4
21.4
0.0
DME
23.5
31.0
0.0
Hydrocarbon
0.2
14.6
44.0
Hydrocarbon
CH4
1.5
1.1
1.1
distribution*,
C2H6
-
0.1
0.6
wt%
C3H8
2.0
2.5
16.2
i-C4H10
13.8
6.5
18.7
n-C4H10
-
1.3
5.6
C2H4
18.1
12.4
0.5
C3H6
48.2
26.7
1.0
C4H8 (butenes)
11.9
15.8
1.3
C5+ aliphatic HC
4.4
27.0
14.0
Aromatic HC
-
6.6
41.1
Aromatic HC
Benzene
4.1
Page 40 of 46
distribution*,
Toluene
25.6
wt%
Ethylbenzene
1.9
(normalized)
Xylenes
41.8
Trimethylbenzenes
14.1
Ethyltoluenes
4.8
Isopropylbenzene
0.2
Tetramethylbenzenes
4.3
Estimated H/C ratio**
2.1
Pressure for liquefaction**,
bar
at -50 °C
5.6
at -100 °C
2.0
Page 41 of 46
General References
Bissett, L. 1977. “Equilibrium Constants for Shift Reactions”. Chemical Engineering, 84
(23):155
Dubois, J. -L., Sayama, K., Arakawa, H. 1992. “CO2 Hydrogenation over Carbide
Catalysts”. Chem. Lett., 5-8.
Ernst, K.-H., Campbell, C., Moretti, G. 1992. “Kinetics of the Reverse Water-Gas Shift
Reaction over Cu(110)”. J. Catal., 134, 66-74.
Ewell, R.H. 1940. “Calculation of Chemical Equilibrium at High Pressures”. Ind. Engr.
Chemistry, 32 (2), 147-153.
Fujita, S., Usui, M., Takezawa, N. 1992. “Mechanism of the Reverse Water Gas Shift
Reaction over Cu/ZnO Catalyst”. J. Catal. 134, 220-225.
Kitayama, Y., Watanabe, Y., Muramatsu, K., Kodama, T. 1997. “Catalytic Reduction of
Carbon Dioxide on Ni-Cu Alloys”. Energy, 22, 177-182.
Nozaki, F., Sodesawa, T., Satoh, S., Kimura, K. 1987. “Hydrogenation of Carbon
Dioxide into Light Hydrocarbons at Atmosheric Pressure over Rh/Nb2O5 or Cu/SiO2Rh/Nb2O5 Catalyst”. J. Catal., 104, 339-346.
Saito, M., Anderson, R. B. 1981. “The Activity of Several Molybdenum Compounds for
the Methnation of CO2”. J. Catal., 67, 296-302.
Sakurai, H., Tsubota, S., Haruta, M. 1993. “Hydrogenation of CO2 over Gold Supported
on Metal Oxides”. Appl. Catal. A., 102, 125-136.
Sakurai, H., Ueda, A., Kobayashi, T., Haruta, M. 1997. “Low-Temperature Water-Gas
Shift Reaction over Gold Deposited on TiO2.” Chem. Commun., 271-272.
Suzuki, T., Iwanami, H. -I., Yoshizawa, T., Yamazaki, H., Yoshida, Y. 1995. “Selective
Hydrogenation of CO2 to CO in the Presences of Excess CO on MoO3/ZnO Catalyst. A
Preliminary Attempt to Yield an Equimolecular Mixture of H2 and CO”. Int. J. Hydrogen
Energy, 20, 823-830.
Taoda, H., Osaki, T., Iseda, K., Horiuchi, T., Yamakita, H. 1991. “High Activity
Supported Molybdenum Sulfide Catalyst for Selective Reduction of CO2 to CO”. Chem.
Express, 6, 1013-1016.
Tingey, G.L. 1966. “Kinetics of the Water-Gas Equilibrium Reaction. I. The Reaction
of Carbon Dioxide with Hydrogen”. J. Phys. Chemistry, 70 (5), 1406-1412.
Page 42 of 46
Thomas, W. J., Portalski, S. 1958. “Thermodynamics in Methanol Synthesis”. Ind.
Engr. Chemistry, 50 (6), 967-970.
DME Synthesis References
Karpuk, M. E. and Cowley, S. W. 1988. “On-Board Dimethyl Ether Generation to Assist
Methanol Engine Cold Starting.” Proceedings of the International Fuels and Lubricants
Meeting and Exposition, SAE Paper 881678.
Sofianos, A. C. and Scurrell, M. S. 1991. “Conversion of Synthesis Gas to Dimethyl
Ether over Bifunctional Catalytic Systems.” Ind. Eng. Chem. Res. 30, 2372-2378.
Li, J.-L., Zhang, X.-G. and Inui, T. 1996. “Improvement in the Catalyst Activity for
Direct Synthesis of Dimethyl Ether from Synthesis Gas through Enhancing Dispersion of
CuO/ZnO/g-Al2O3 in Hybrid Catalysts.” Appl. Catal. A, 147, 23-33.
F-T Synthesis References
Snel, R. 1988. “Catalytic Hydrogenation of Carbon Monoxide to Alkenes Over Partially
Degraded Iron Complexes. I. Unsupported Iron Catalysts.” Appl. Catal. A, 37, 35-44.
Snel, R. 1989. “Catalytic Hydrogenation of Carbon Monoxide to Alkenes Over Partially
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Huang, C.-S., Dabbagh, H. A. and Davis, B. H. 1991. “Fisher-Tropsch Synthesis. A
Measure of the Contribution of Hydrogenolysis Using a Doubly Promoted Iron Catalyst
in a Continuous Stirred Tank Reactor.” Appl. Catal., 73, 237-248.
Zimmerman, W. H., Rossin, J. A. and Bukur, D. B. 1989. “Effect of Particle Size on the
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Page 43 of 46
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1989. “Binder/Support Effects on the Activity and Selectivity of Iron Catalysts in the
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MTG Process References
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Chang, C. D. and Silverstri, A. J. 1977. J. Catal., 47, 249.
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Page 44 of 46
Form Approved
OMB No. 0704-0188
REPORT DOCUMENTATION PAGE
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1. AGENCY USE ONLY (Leave blank)
4.
2. REPORT DATE
September 17, 1997
3. REPORT TYPE AND DATES COVERED
SBIR Final Report; 03/17/97 – 09/17/97
TITLE AND SUBTITLE
Methanol Mars In-Situ Propellant Production – Final Report
5.
FUNDING NUMBERS
Contract No.
6.
AUTHORS
Robert M. Zubrin, Brian Frankie, and Tomoko Kito
7.
PERFORMING ORGANIZATION NAME(S) AND ADDRESS(ES)
NAS 9-97082
8.
PERFORMING ORGANIZATION
REPORT NUMBER
Pioneer Astronautics
445 Union Blvd., Suite 125
Lakewood, CO 80228
9.
PA-ISRU-1
SPONSORING/MONITORING AGENCY NAME(S) AND ADDRESS(ES)
10. SPONSORING/MONITORING AGENCY
REPORT NUMBER
NASA Lyndon B. Johnson Space Center
2101 NASA Road 1
Houston, TX 77058
11. SUPPLEMENTARY NOTES
12a. DISTRIBUTION/AVAILABILITY STATEMENT
12b. DISTRIBUTION CODE
For general distrbution
13. ABSTRACT (Maximum 200 words)
The Methanol Mars In-Situ Propellant Production (MMISPP) Project built a unit for producing both storable fuel and oxygen on
the surface of Mars with more than 95% of the required raw material mass derived from the Martian atmosphere. In the MMISPP
system, a reverse water gas shift reactor combines a small quantity of imported hydrogen with Martian atmospheric CO2 to
produce a syngas feed for a catalytic methanol or other fuel synthesis reactor. Water from the RWGS can be electrolyzed to
produce oxygen and return hydrogen feedstock to the system. The synthesized fuel/oxygen bi-propellant produced can be used as
either rocket propellant or fuel for ground vehicles. The project demonstrated production of methanol as well as dimethyl ether and
higher hydrocarbons. The project results show an attractively high leverage ratio, the flexibility to produce a number of different
type of fuels, and a power consumption about half that of most alternative Mars in-situ propellant production systems currently
being researched.
14. SUBJECT TERMS:
15. NUMBER OF PAGES
45
Mars Methanol In Situ Resource Utilization; Methanol Production; Oxygen Production;
Propellant Production; Reverse Water Gas Shift Reaction
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Unclassified
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