ANALYSIS OF MIXING EFFECTS ON THE METHANOL TO OLEFINS REACTION

By

CHRISTOPHER MICHAEL FISCHER

A thesis submitted in partial fulfillment of the requirements for the degree of

MASTER OF SCIENCE IN CHEMICAL ENGINEERING

WASHINGTON STATE UNIVERSITY Department of Chemical Engineering

December 2001

ii

To the Faculty of Washington State University:

The members of the Committee appointed to examine the thesis of CHRISTOPHER

MICHAEL FISCHER find it satisfactory and that it be accepted.

_________________________________ Chair _________________________________ _________________________________

iii

ACKNOWLEDGMENTS

Through the course of completing my degree at the Masters of Science level I have had

the good fortune of acquiring the help and guidance of many individuals whom I would like to

give recognition. I would first like to thank Dr. William Thomson for advising my through both

my undergraduate and graduate studies and passing on to me his knowledge in the field of

catalysis. I would also like to thank the members of my committee, Dr. Reid Miller and Dr.

Richard Zollars, both whom I have learned from their instruction in various chemical

engineering classes. Dan Leatzow, I would like to thank you for all the advice and guidance you

have given to me, without which could have made the completion of this degree more difficult.

Finally, I would like to give special thanks to all my family and friends. Your support and

encouragement over the years has kept me strong and allowed me to push on when things

seemed impossible.

iv

ANALYSIS OF MIXING EFFECTS ON THE METHANOL TO OLEFINS REACTION

Abstract

by Christopher Michael Fischer, M.S.

Washington State University

December 2001

Chair: William J. Thomson

An investigation was conducted to examine the effect of reactor mixing states on the

methanol-to-olefins (MTO) product distribution when using a ZSM-5 catalyst. A plug flow

reactor (PFR) and a continuous stirred tank reactor (CSTR) were used to produce a plug flow and

a backmix mixing state under reaction conditions of 450ºC and 8 psig. In order to compare the

effect of mixing states at the same conversion and coke levels, spent catalysts from the CSTR

were used in the PFR to obtain initial selectivity data. At coke levels from 1.39% to 2.60% and

conversions between 36 and 82%, the only direct effects of mixing states were an increase in the

saturates selectivity and lack of carbon monoxide in the CSTR. However, the type of coke

deposited on the catalysts was dependant on the mixing state and influences the

ethylene:propylene ratio at conversions below 80%. The results also indicate that there may be

three parallel methanol reactions under these conditions: (1) direct conversion to olefins, (2)

conversion to olefins via the formation of dimethyl ether, and (3) decomposition of methanol to

carbon monoxide and hydrogen. The decomposition reaction was far more prevalent in the PFR

and thus, occurs at higher methanol partial pressures.

v

ACKNOWLEDGEMENTS……………………………………………..……………………… iii

ABSTRACT………………………………………………………………………………………iv

LIST OF FIGURES…………...…………………………………………………………………vii

LIST OF TABLES……………………………………………………………………………… xii

CHAPTER

1. INTRODUCTION……………………………………………………………………… 1

2. BACKGROUND……………………………………………………………………… 3

Establishment of the methanol to olefins process…………………………………… 3

Properties of ZSM-5………………...……………………………………………… 7

Effects of space velocity………………………………………………………………9

Effects of Si/Al ratio………………………………………………………………… 10

Effects of temperature…………………………………………………………………11

Effects of partial pressure………………………………………..……………………14

Deactivation and coke formation…………………………………………………… 16

Effects of mixing states……………………………………………………………… 19

3. EQUIPMENT AND METHODOLOGY………….…...……………………………… 21

Equipment…………………………………………………………………………… 22

A. Apparatus……………...……………………………………………………22

B. Berty reactor……………..…………………………………………………25

C. Plug flow reactor………………………..………………………………… 32

Methodology………………………………………………………………………… 32

TABLE OF CONTENTS

vi

A. Catalyst preparation………………………….…………………………… 36

B. Initial equipment/catalyst testing……………………………..…………… 37

C. Determination of coke levels…………………………...…...………………37

D. Base cases for the PFR and Berty Reactor…………………………………38

E. Berty reactor experiments……………………...……………………………38

F. PFR experiments…………………..…………….………………………… 39

4. RESULTS AND DISCUSSION…………………………...……………………………41

Initial equipment/catalyst testing………………………………………………………41

A. Equipment testing……………………...……………………………………41

B. Catalyst testing……………………...………………………………………44

Base cases for the PFR and Berty reactors………………………………….……… 58

Berty reactor experiments…………………………………………………………… 69

PFR experiments………………………………………………………………………77

Comparison of product distributions………………………………………………… 83

5. CONCLUSIONS AND RECOMMENDATIONS…………………………..………… 87

REFERENCES………………………………………………………………….……………… 90

APPENDIX………………………………………………………………………………………97

A. GCD Calibration………………………….…………………………………………… 98

B. Reactor(s) operations procedures…………………..………………………………… 100

C. Procedure for preparing and operating the TGA/MS equipment. ………………………109

D. Catalog of data collected for MTO investigation………………………………………114

E. Carbon balance for PFR Run # 11 [ZSM-5 (L)] (Normalized without Water)…………132

vii

Figure 1: Schematic of the hydrocarbon pool mechanism…………………………………………6

Figure 2: Simplified sketch of ZSM-5 internal pore structure…………………………………… 8

Figure 3: Schematic of experimental setup used for the MTO mixing state investigation.

The apparatus consisted of three main parts: 1-vaporizer system, 2-reactor, 3-

backpressure regulating system. Points labeled A through J indicate the locations of thermocouples used to monitor the temperature of the heat tapes. All lines

that made contact with the reactant feed and product stream required heating

to prevent any condensation. ……………………………………………………………… 23

Figure 4: Photograph of the Berty reactor……………………………………………………… 26

Figure 5: Schematic of the Berty reactor. Note the flow patterns produced for the back

mixing environment………………………………………………………………………… 27

Figure 6: Close up of the magnetic drive system for the Berty reactor. Note the cooling

lines connected to a cooling jacket located just above the magnetic assembly..…………… 28

Figure 7: Close up of the catalyst basket unassembled (top) and assembled (bottom).

This basket rests in place on the lower half of the Berty reactor just about the

mixing impeller..………………………………….………………………………………… 29

Figure 8: Copper gaskets and stainless steel spacer used to seal Berty reactor (top).

Upper copper gasket and stainless steel spacer in place on the upper half of the

Berty reactor (bottom).………………………………………………………………………30

Figure 9: Close up of heating jacket used for the Berty reactor………………………………… 31

Figure 10: Photo of PFR secured to the upstream and downstream bulkheads. The

LIST OF FIGURES

viii

thermocouple used to measure the temperate of the catalyst bed was inserted

through the exit end of the reactor……………………………………………………………33

Figure 11: Stainless steel wire mesh tube that was inserted into the exit end of the

reactor to hold the catalyst in place (the mesh tube is open only at one end). The

catalyst was loaded inside of the wire mesh tube while the quartz beads rested on

top of the mesh tube (on top of the closed end)…………………………………………… 34

Figure 12: Flowchart of methodology for MTO mixing state investigation.……………………. 35

Figure 13: Mole fraction of ethylene versus time on stream for 5.00, 9.03, and 15.05

WHSV using 0.50 g ZSM-5 (initial batch). Reactor at 450ºC and 8 psig, sample

valve at 2 psig, 85% helium dilution. [Water, dimethyl ether, methanol free basis]…………47

Figure 14: Mole fraction of propylene versus time on stream for 5.00, 9.03, and 15.05

WHSV using 0.50 g ZSM-5 (initial batch). Reactor at 450ºC and 8 psig, sample

valve at 2 psig, 85% helium dilution. [Water, dimethyl ether, methanol free basis]…………48

Figure 15: Mole fraction of butene versus time on stream for 5.00, 9.03, and 15.05

WHSV using 0.50 g ZSM-5 (initial batch). Reactor at 450ºC and 8 psig, sample

valve at 2 psig, 85% helium dilution. [Water, dimethyl ether, methanol free basis]…………49

Figure 16: Mole fraction of saturates versus time on stream for 5.00, 9.03, and 15.05

WHSV using 0.50 g ZSM-5 (initial batch). Reactor at 450ºC and 8 psig, sample

valve at 2 psig, 85% helium dilution. [Water, dimethyl ether, methanol free basis]…………50

Figure 17: Mole fraction of carbon monoxide versus time on stream for 5.00, 9.03, and

15.05 WHSV using 0.50 g ZSM-5 (initial batch). Reactor at 450ºC and 8 psig,

sample valve at 2 psig, 85% helium dilution. [Water, dimethyl ether, methanol free

basis]……...………………..…………………………………………………………………51

ix

Figure 18: Mole fraction of dimethyl ether versus time on stream for 5.00, 9.03, and

15.05 WHSV using 0.50 g ZSM-5 (initial batch). Reactor at 450ºC and 8 psig,

sample valve at 2 psig, 85% helium dilution. [Water free basis]……………………………53

Figure 19: Mole fraction of methanol ether versus time on stream for 5.00, 9.03, and

15.05 WHSV using 0.50 g ZSM-5 (initial batch). Reactor at 450ºC and 8 psig,

sample valve at 2 psig, 85% helium dilution. [Water free basis]……………………………54

Figure 20: Average mole fraction of light olefins and saturates versus time on stream for

the three PFR base case runs. Reactor at 450ºC and 8 psig, sample valve at 2 psig,

85% total dilution, 9.03 WHSV, 0.50 g ZSM-5 (master batch). [Water, argon,

dimethyl ether, methanol free basis]……………………..……………………………………60

Figure 21: Average mole fraction of carbon monoxide and carbon dioxide versus time on

stream for the three PFR base case runs. Reactor at 450ºC and 8 psig, sample valve

at 2 psig, 85% total dilution, 9.03 WHSV, 0.50 g ZSM-5 (master batch).

[Water, argon, dimethyl ether, methanol free basis]……………………..………………… 61

Figure 22: Average mole fraction of methanol and dimethyl ether with total oxygenates

conversion versus time on stream for the three PFR base case runs (averaged).

Reactor at 450ºC and 8 psig, sample valve at 2 psig, 85% total dilution, 9.03 WHSV,

0.50 g ZSM-5 (master batch). [Water, argon free basis]……………………………………62

Figure 23: Average mole fraction of light olefins and saturates versus time on stream for

the three Berty reactor base case runs. Reactor at 450ºC and 8 psig, sample valve at

2 psig, 80% total dilution, 0.903 WHSV, 10.0 g ZSM-5 (master batch). [Water,

argon, dimethyl ether, methanol free basis]………………………………………………… 65

Figure 24: Average mole fraction of carbon dioxide versus time on stream for the three

x

Berty reactor base case runs. Reactor at 450ºC and 8 psig, sample valve at 2 psig,

80% total dilution, 0.903 WHSV, 10.0 g ZSM-5 (master batch). [Water, argon,

dimethyl ether, methanol free basis]…………………………………………………………66

Figure 25: Average mole fraction of methanol and dimethyl ether with total oxygenates

conversion versus time on stream for the three Berty reactor base case runs.

Reactor at 450ºC and 8 psig, sample valve at 2 psig, 80% total dilution, 0.903 WHSV,

10.0 g ZSM-5 (master batch). [Water, argon free basis]……………………………………67

Figure 26: Example of a TGA/MS analysis (From Run 10) The temperature of the

furnace was held at 200ºC for removal of water, then held at 800ºC for coke

removal. The plot contains both the TGA data plus the MS data for carbon

dioxide. The CO2 peak was used to indicate the beginning and end of coke

removal. The first mass change is water removal; the second mass change is

coke removal. (Feed is 8.8% O2)……………………………………………………………72

Figure 27: Mole fraction of light olefins and saturates for the four coked master

batches of catalyst from the Berty reactor. Reactor at 450ºC and 8 psig, sample

valve at 2 psig, 80% total dilution. Coke levels are 1.39, 1.41, 2.60 and 2.40 wt%

for 81.8, 36.1, 69.6, and 38.2% oxygenate conversion, respectively. [Water,

argon, dimethyl ether, methanol free basis]………………………………………….………73

Figure 28: Mole fraction of methanol and dimethyl ether for the four coked master

batches of catalyst from the Berty reactor. Reactor at 450ºC and 8 psig, sample

valve at 2 psig, 80% total dilution. Coke levels are 1.39, 1.41, 2.60, and 2.40 wt%

for 81.8, 36.1, 69.6, and 38.2% oxygenate conversion, respectively. [Water,

argon free basis]…………………………………………………………………….……… 74

xi

Figure 29: Schematic of the corrected hydrocarbon pool mechanism……………………………76

Figure 30: Mole fraction of light olefins and saturates for PFR runs using the 1.39 wt%

coked catalyst from the Berty reactor. Reactor at 450ºC and 8 psig, sample valve

at 2 psig, 80% total dilution, 0.25 g ZSM-5. The data points at 70% oxygenates

conversion are averages of PFR runs 3-5. [Water, argon, dimethyl ether,

methanol free basis]…………………………………………………………….……………80

Figure 31: Mole fraction of light olefins and saturates for PFR runs using the 2.60 wt%

coked catalyst from the Berty reactor. Reactor at 450ºC and 8 psig, sample valve

at 2 psig, 80% total dilution, 0.25 g ZSM-5. The data points at 62.1% oxygenates

conversion are averages of PFR runs 2-4. [Water, argon, dimethyl ether,

methanol free basis]…………………………………………………………………………81

Figure 32: Ethylene and propylene mol fraction comparison between the coked Berty

catalyst and fresh catalyst (Fresh catalyst data is from Run 8). Conversions for

the 1.39 and 2.60 wt% coke are 81.4 and 81.3% (above) and 70.0 and 62.1%

(below). Conversion was not measured for Run 8. [Water, argon, dimethyl ether,

methanol free basis]…………………………………………………………………………82

Figure 33: Mole fraction of light olefins and saturates for the PFR (81.4% conversion,

1.46 wt% coke) and Berty reactor (81.8% conversion, 1.39 wt% coke).

[Water, argon, dimethyl ether, methanol free basis]…………………………………………84

Figure 34: Mole fraction of light olefins and saturates for the PFR (62.1% conversion,

2.63 wt% coke) and Berty reactor (69.6% conversion, 2.60 wt% coke).

[Water, argon, dimethyl ether, methanol free basis]…………………………………………85

xii

Table 1: Flow parameters and results for PFR mixing tests (blank, no catalyst)…………………43

Table 2: Flow parameters used for examining the effects of space velocity on the MTO

product distribution. Reactor at 450ºC, 8 psig, sample valve at 2 psig, 85%

dilution, 0.50 g ZSM-5 (initial batch)……………………………………………………… 46

Table 3: Flow parameters and acquired methanol calibration constants for recalibration

of methanol at 85% helium dilution (Blank run, no catalyst)…………………………………56

Table 4: Experimental parameters used for base case studies (PFR and Berty reactor)…………59

Table 5: Product comparison between the initial (Run 8) and base case PFR runs

for 9.03 WHSV. Reactors at 450ºC, 8 psig, sample valve at 2 psig, 85% dilution,

0.50 g ZSM-5 used in each data series. [Water, argon, dimethyl ether, methanol

free basis]……………………………………………………………………………………63

Table 6: Product comparison between the PFR and Berty base case runs. Reactors at

450ºC, 8 psig, sample valve at 2 psig. PFR loaded with 0.50 g ZSM-5 (master

batch) with 85% feed dilution. Berty reactor loaded with 10.0 g ZSM-5 (master

batch) with 80% feed dilution. [Water, argon, dimethyl ether, methanol free

basis except for dimethyl ether and % oxygenates conversion (these are on a

water, argon free basis)]…………………………………………………………………… 68

Table 7: Experimental parameters and measured coke content of Berty runs that will

be used in comparison PFR experiments……………………………………………………70

Table 8: Experimental parameters and measured conversions for PFR runs that used

coked catalyst from the Berty reactor. Final conversions measured at 80%

LIST OF TABLES

xiii

dilution with 10.0 sccm argon in mixture. All runs used 0.25 g of catalyst per

loading. Time on stream for each run before sampling was 1 minute.………………………78

1

CHAPTER 1

INTRODUCTION

The petrochemical industry produces many products that civilization has come to use and

depend on. However, as time progresses, traditional components used to make these products

begin to decrease in availability, which drives the research community to pursue new raw

materials and techniques that have the potential to replace traditional petrochemical industrial

practices. For example, key components such as light olefins (C2=-C4

=) are traditionally

produced by thermal cracking of naphtha. With increasing prices and decreasing quantities of

oil, an alternative means of producing these light olefins would prove invaluable. A possible

means of producing the light olefins is by the direct conversion of methanol, which has become

known as the methanol-to-olefins reaction, or MTO. The O.H. Reaugh lab at Washington State

University has been investigating the MTO reaction, or more specifically, the effects that mixing

states have on the MTO product distribution. Little work has been published that focuses on the

effects of reactor mixing states on the MTO product distribution. Knowledge in this area could

prove beneficial for establishing and improving the operation of industrial scale MTO facilities.

The format of this thesis consists of an overview of previous work conducted for the MTO

reaction, a detailed discussion of the equipment and methodology employed in the investigation,

and a discussion of the results and closing conclusions. Appendices are included to give detailed

outlines of how various pieces of equipment were calibrated and the detailed methodology used

to acquire the data for the investigation.

Two reactors were used to provide the different mixing states: a plug flow reactor and a

Berty reactor (continuous stirred tank reactor). Experiments were conducted with each reactor at

2

the same reaction conditions (temperature, partial pressure of methanol in the feed, coke levels,

and conversion of reactants). The product streams from each reactor were measured for

composition and used to examine the effects of the different mixing states.

3

CHAPTER 2

BACKGROUND

Establishment of the methanol-to-olefins process

Molecular sieve catalysts have become an efficient means to selectively produce desired

products while minimizing the production of undesired byproducts. These catalysts contain a

distribution of acid sites, which are ideal for hydrocarbon reactions and have pore sizes that are

on the molecular scale, allowing only molecules of a specific size range to enter and exit the

catalyst pores. An example of applying a molecular sieve catalyst is Mobil’s methanol-to-

gasoline process (MTG). Mobil found that by using a zeolite catalyst called ZSM-5, methanol

could be converted directly into high-octane gasoline [1]. Methanol was chosen due to the

availability of natural gas, which can be easily converted to methanol. During the MTG process,

light hydrocarbon olefins are formed as intermediates before the ZSM-5 oligomerizes them to

higher molecular-weight olefins (components of gasoline, C5-C11’s). This process can be

generalized by the following equation [2], [3], [4]:

olefinsC aromatics

paraffins-isoCCOCHCHOHCH2

6

52OH

33OH

322

+

==−± →− → →← (1)

The first step in Equation 1 is the dehydration of methanol to dimethyl ether (DME) and water.

Typically, when the methanol conversion value is reported, it also includes the conversion of

DME. This is because DME is the species that actually forms the hydrocarbons after the initial

dehydration step. A better way of reporting reactant conversion is by indicating, “oxygenate”

4

conversion, which includes both methanol and DME. The resulting equilibrium mixture from

the initial dehydration shown in Equation 1 then forms the light end olefins, C2=-C5

=.

These olefins, particularly the C2=-C4

= olefins, happen to be useful intermediates in the

petrochemical industry [5], especially ethylene and propylene. It is estimated that 70% of

domestic ethylene production is used as feedstock to produce downstream products such as

polyethylene, ethylene dichloride, ethylene glycol, ethylbenzene, linear alcohols, and vinyl

acetate [6]. Propylene is used to produce products such as polypropylene, acrylonitrile,

propylene oxide, cumene, oxo-alcohols, isopropanol, oligomers, and acrylic acid [7]. Due to

predictions that the markets for ethylene and propylene will increase considerably in the future, a

strong emphasis has been placed on finding a more efficient and easier means of converting

methanol to light end olefins [8]. Whereas the MTG process is typically operated at 400°C and a

pressure of several bars, it was observed that if the operating conditions for the MTG process are

altered while using ZSM-5 (such as increasing the temperature from 400°C to 500°C), the light

end olefins can be harvested before completion of the MTG process, hence the beginning of the

MTO process [1]. In 1979, Mobil demonstrated the potential for commercialization of the MTO

process in an experimental plant in Wesseling (Germany). This encouraged Universal Oil

Products (UOP) to experiment with medium and small pore zeolites targeted for olefin

extraction. The result was SAPO-34 (silicoaluminophosphates), a small pore zeolite. SAPO-34

demonstrated high selectivity towards ethylene and propylene, and also proved to be flexible,

since the ratios of the two olefins could be adjusted by changing the operating conditions of the

reaction and this was demonstrated and confirmed by Norsk Hydro using a 0.5 ton-year-1 pilot

plant [1].

There is a large debate on the correct MTO mechanism for production of primary

5

products, that is, where the first C-C bond is formed from the methanol/DME reactants. There

are many theories presented in the literature (more than 20). Out of all of the theories, the

“parallel” mechanism appears to be the most favorable.

The ‘parallel’ mechanisms for the MTO process consist mainly of what is called the

‘hydrocarbon pool’. In 1994, Dahl and Kolboe [9], [10] suggested the following scheme for the

hydrocarbon pool shown in Figure 1. The (CH)n in Figure 1 represents an adsorbate that may

have many characteristics in common with ordinary coke, and which might easily contain less

hydrogen than indicated. It could possibly be represented better by (CHx)n with 0<x<2. Dahl

and Kolboe formulated Figure 1 by conducting isotopic labeling experiments where

[12C]ethylene and [12C]propylene were run with [13C]MeOH over 100 mg of SAPO-34 in a fixed

bed reactor. Essentially, they observed that only a minor fraction of the propylene molecules

were formed by the addition of methanol to ethylene (propylene molecules formed from the

addition of methanol are identified by a 12C/13C ratio larger than one). The majority of the

propylene molecules were formed directly from methanol, supporting the presence of a

‘hydrocarbon pool’.

Dahl and Kolboe’s results do not address the nature of the hydrocarbon pool mechanism.

The ‘rake’ mechanism is actually an earlier version of the ‘pool’ mechanism that originates from

Cormerais et al. in 1980 [11]. According to Cormerias et al., the hydrocarbon pool consists of

ethoxy, propoxy, butoxy groups, etc., formed by methylation of the next lower member. The

authors concluded from their results that the reaction proceeds via an adsorbate that is

continually adding and splitting off reactants and products. These results were in agreement with

a ‘hydrocarbon pool’ mechanism, and ethylene was observed as an inert primary product (Dahl

and Kolboe also observed that ethylene was an inert product). In the co-reaction of ethylene and

6

Figure 1: Schematic of the hydrocarbon pool mechanism.

7

methanol over SAPO-34, most of the product was made from methanol, and there was about

90% unconverted ethylene remaining upon 100% conversion of methanol. Exiting propylene

and butene formed during the reaction showed distributions of 12C and 13C, which are nearly

random at all times. This observation again supports the ‘hydrocarbon pool’ theory, although,

randomness will also result if propylene and butene were adsorbed, isotopically isomerized, and

desorbed at a much faster rate than ethylene [9], [10].

Properties of ZSM-5

SAPO-34 and its different isotopes have been found to achieve the best olefin selectivites

when compared to other zeolites. However, the catalyst is not readily available, and synthesis of

SAPO-34 is very complex, so it could not be obtained for this investigation. ZSM-5, which is

readily available from several commercial suppliers, can be used as an alternative to run the

MTO reaction and was substituted in place of SAPO-34. ZSM-5 is a medium pore zeolite with a

tetrahedrally coordinated zeolite lattice that contains aluminum cations. Medium pore zeolites

have pore sizes that range between 0.5 to 0.6 nm. The following empirical formula can be used

to represent the composition of ZSM-5 [2], [12]:

3abouttypicallyand27nwithOHOSiAlNa 2192n96nn <⋅− (2)

Different isotopes of ZSM-5 can also be synthesized by isomorphously substituting other metal

cations for the aluminum cations. Two perpendicularly intersecting channel systems make up

the ZSM-5 structure as shown in Figure 2 [2]. The actual active sites on the catalyst are believed

to be located at the intersections produced by the parallel channels. These intersections are

8

Figure

Location of an active site

2: Simplified sketch of ZSM-5 internal pore structure.

9

approximately 0.9 nm in diameter [2], [13].

Zeolites such as ZSM-5 typically have strong acid and carboniogenic properties [2], [14].

Tetravalent silica is substituted by trivalent aluminum, providing the acidic characteristics for the

zeolite. Although the total acidity is a combination of Lewis and Brönsted acid sites, the Lewis

sites are believed to contribute little to the conversion of hydrocarbons. Anderson et al.

demonstrated that zeolites’ Brönsted acid sites are involved with the conversion of methanol, and

the water produced in the reaction would quench any Lewis acidity [15], [16]. However,

Wolthuizen et al. utilized ethylene chemisorption to show that Lewis acid sites enhance the

polymerization of ethylene [17].

The main advantage to using zeolites for MTO is that the small pore sizes provide a

sieving effect that enhances the selectivity towards the light end olefins. Larger molecules are

not allowed to enter or exit the catalyst pores, restricting passage to the smaller species such as

ethylene and propylene. Unfortunately, the small pore sizes that provide excellent sieving

properties also causes difficulty with quick catalytic deactivation due to blocking of the acid sites

from coke formation. This is especially a problem with the small pore zeolites such as SAPO-34

(the pores of SAPO-34 are only 0.45 nm). However, investigations have shown that the shape-

selectivity of the medium pore zeolites, especially ZSM-5, actually increases the resistance to

deactivation by coking [18], [19], [20].

Effects of space velocity

Studies have shown that using different space velocities can change the product

distribution for MTO when using ZSM-5. Chang et al. observed through a series of space

velocity experiments that the C2=-C5

= olefins are intermediates as methanol is converted to

10

hydrocarbons [21]. While maximum olefin selectivites were observed at 90% conversion, only

45.7 wt% of the hydrocarbon products were olefins. Another point to mention is that when

100% conversion of methanol was achieved, the ZSM-5 catalyst continued to convert the smaller

hydrocarbons (meaning the desired light olefins) to larger hydrocarbon species. Consequently,

caution should be used when examining MTO product distributions when running at complete

oxygenate conversion.

Space velocity experiments have also indicated that the MTO reaction may be

autocatalytic where the alkylation of light olefins with methanol occurs. In support of the

autocatalytic effect, Chen et al. observed that the conversion of methanol or DME to

hydrocarbons was very slow at high space velocities (low conversion) [22]. Using data acquired

from an isothermal reaction system, they observed that the rate of higher olefin formation via an

autocatalytic route was 50 times faster than the rate of initial ethylene formation from

methanol/DME. However, in a narrow range of space velocities, the rate of methanol or DME

conversion accelerates rapidly with increasing concentration of hydrocarbons. Espinoza

disagreed with the autocatalytic interpretation, indicating that the sudden increase in methanol

conversion to hydrocarbons by changing the space velocity was more likely an effect of meeting

a required concentration of the intermediate DME [2], [23]. By feeding DME as a feed in his

investigation, the sudden jump in conversion was not observed.

Effects of Si/Al ratio

The Si/Al ratio of ZSM-5 affects the amount of observed catalytic activity [24].

Increasing the Si/Al ratio decreases the acidity of ZSM-5, hence decreasing the activity of the

catalyst. However, as the activity decreases, the selectivity for the light end olefins increases.

11

Studies performed by Dehertog et al. showed that maximum olefin selectivity is achieved at

higher methanol conversions as the Si/Al ratio is increased [25]. The maximum olefin selectivity

for these studies was achieved at Si/Al = 200. Further increase in Si/Al ratio resulted in a

decrease in the maximum achievable olefin selectivity. These results agree with the observations

of Chang et al. who achieved the highest olefin yields at Si/Al = 250 using HZSM-5 (HZSM-5 is

the hydrogen form of ZSM-5) [21]. Prinz and Riekert achieved an optimized maximum

selectivity with Si/Al = 100 using HZSM-5 [26].

Wu and Kaeding also reported that lower catalytic activity produces higher olefin

selectivites with ZSM-5 using Si/Al ratios ranging from 35 to 1600 [27]. In their investigation,

Wu and Kaeding used the α-value as a measure of catalyst activity (the α-value was an arbitrary

unit created by Weisz and Miale to measure the relative activity of a catalyst) [28]. Large α-

values indicated high catalytic activity; small α-values indicate low catalytic activity. Their

results demonstrated at complete methanol conversions that the catalysts used (referring mainly

to ZSM-5) must have a α−value lower than 30 to obtain high selectivites for light olefins.

Effects of temperature

The effects of temperature on the MTO product distribution when using ZSM-5 catalysts

are similar to those observed by varying the space velocity at constant temperature [3]. Chang et

al. investigated the affects of temperature on the MTO product distribution in the range of 400-

500°C using HZSM-5 (Si/Al ratio = 250) [21], and hypothesized the following kinetic model for

their investigation:

CBA 21 kk →→ (3)

12

In equation (3), “A” represents the oxygenates (as CH2), “B” is the light end olefin products, and

“C” is the aromatics and paraffins. The activation energy for olefin formation was found to be

19.3 kcal/mol while there was little temperature dependence for aromatic formation. This led to

the conclusion that olefin formation is favored over aromatics formation at higher temperatures.

It was also observed that the ratio k1/k2 (a measure of selectivity) was dependant on the Si/Al

ratio. Increasing the Si/Al ratio increased the k1/k2 ratio, which demonstrated that there was a

difference in dependency on the Brönsted acidity between the two reactions. From these

observations, the conclusion was made that olefin production can be improved by using low

catalyst acidity and high operating temperatures.

In terms of thermodynamics, increasing the temperature over a similar range of olefin

partial pressures shifts the products to the lower olefins [29]. However, if it is assumed that the

only reaction in the system is the formation of olefins, this equilibrium product distribution can

only be approached at low methanol conversions. Low methanol conversions are necessary for

suppressing any further reaction of the olefin products to aromatics and saturates. For example,

Chu and Chang reported ethylene selectivites that were ~100% at methanol conversions less than

2% at 500ºC. As methanol conversion is increased, kinetics has a stronger influence in

governing the product distribution due to autocatalysis. It is believed that the autocatalytic step

involves the alkylation of light olefins with methanol [2]. Due to this kinetic observation, the

C3+ olefin selectivites will exceed those calculated by thermodynamics.

In an investigation conducted by Dehertog and Froment with HZSM-5, a strong

temperature dependence of the olefin yield and distribution was observed [25]. The light end

olefin yield was nearly doubled from 14.4 to 27.0 g/100 g methanol fed by increasing the

temperature from 320 to 480ºC. This temperature increase also increased the fraction of C2-C4

13

olefins in the total hydrocarbon distribution from 33.5 to 62.9 wt%. Dehertog and Froment also

observed that ethylene was the most abundant light end olefin at lower temperatures, but at

higher temperatures propylene and butene became more pronounced in concentration.

An interesting observation reported by Ono et al. was an abrupt jump in the conversion of

oxygenates occurred as the temperature increased from 280 to 300ºC at constant space time over

HZSM-5 [2]. The apparent activation energy associated with the jump is too high for ordinary

chemical reactions, and the “jump” was also observed when feeding DME. This was attributed

to autocatalytic effects in the transformation of dimethyl ether into hydrocarbons. As could be

expected, the temperature at which the “jump” occurred was dependant on the Si/Al ratio. It is

also interesting to note that Espinoza did not observe this conversion “jump” when converting

DME on amorphous, wide-pore acid catalysts [23]. Espinoza concluded that the conversion

“jump” observed with ZSM-5 was due to the onset of cracking of oligomers that form from small

olefins and remain in the medium-sized pores at low temperatures. The occurrence of a “jump”

was further confirmed by similar investigations where propylene was fed over HZSM-5 at

increasing temperatures [2].

Jansen van Rensburg et al. demonstrated that cracking C5+ hydrocarbons at low

temperatures resulted in trace amounts of ethylene, whereas, low conversions of methanol could

produce greater amounts of ethylene [30]. From these experiments, they concluded that ethylene

was a primary reaction product at low conversion and/or low temperature (250ºC) with very little

produced from hydrocarbon cracking. Ethylene then becomes a secondary product from

cracking at higher reaction temperatures (400ºC). Dessau demonstrated that ethylene was not a

primary cracking product when heptene was cracked over HZSM-5 under low pressures [31].

Instead, ethylene was formed by secondary re-equilibration of primary olefin products.

14

Effects of partial pressure

The partial pressure of methanol and dimethyl ether has a significant effect on MTO

olefin selectivity. For example, Chang et al. reported that reduced methanol partial pressures at

370ºC over ZSM-5 suppressed the aromatic reactions, which increased olefin selectivity [32]. In

another investigation by the same authors, a variation of methanol partial pressure from 1.00 bar

to 0.07 bar (constant space time of 70 gcat-h/mole) produced an 8-fold increase in selectivity for

the C2-C4 olefins [2], [33]. This increased the light end olefin fractions from 10.2 to 78.2 wt%.

It is important to mention that not only does lowering the partial pressure increase the light end

olefin selectivity, but it also decreases the reactivity of the system because the rate of reaction is

proportional to the partial pressure of methanol.

To obtain a better representation of the effect of partial pressure on olefin selectivity,

Chang et al. also decreased the space time proportionally with the increase of methanol partial

pressure to maintain a constant reactant contact time [21]. The results indicated a decrease in

olefin selectivity with increasing methanol partial pressure despite the equivalent space times.

Dehertog and Froment examined the increase of methanol space time with decreasing partial

pressures by maintaining complete methanol conversion with space times that ranged from 3.0 to

15.2 gcat-h/mol over HZSM-5 [25]. The yield of light olefins increased by 22.9% from 27.9 to

34.3 g/100 g MeOH fed when the partial pressure was reduced from 1.04 to 0.05 bar. They also

observed that partial pressures have the most pronounced influence on light olefin selectivity at

moderate temperatures of approximately 380ºC. At low temperatures (320ºC) the effects on light

olefin selectivity are only observed at very low partial pressures and are only moderate at higher

temperatures (480ºC).

Dessau used very low partial pressures (1 to 10 Torr) to limit the effects of olefin

15

re-equilibration and examine the true kinetic products of the MTO reaction with HZSM-5 [31].

At high space times and complete methanol conversion, the weight ratio of ethylene to propylene

was less than 0.04. However, at lower space times, ethylene to propylene ratios of 1.0 were

obtained, leading to the conclusion that ethylene is not the initial olefin produced from methanol.

Instead, ethylene is formed by secondary re-equilibrium of primary olefins, propylene, and

butanes. It was also concluded that the formation of ethylene is thermodynamically controlled at

higher partial pressures, even at low conversions. Haag et al. reported contradicting results to

those reported by Dessau with high-silica HZSM-5 (Si/Al = 800) at atmospheric pressure and

low methanol conversions to avoid diffusion/desorption disguises [34]. They concluded that

ethylene was the primary light olefin product, which suggests that propylene and butene are not

primary products.

Caesar and Morrison found that ethylene selectivity could be significantly increased by

reducing the methanol partial pressure with water when using HZSM-5 (315 to 400ºC) [35]. In

support of this observation, Tabak reported increased selectivity toward ethylene when water

was co-fed with methanol or DME [36] and Prinz and Riekert reported that C2 and C3 olefin

selectivity increased along with the ethylene to propylene ratio by adding water to the oxygenate

feed over HZSM-5 [26]. However, Dehertog and Froment observed no effect of water dilution

on light olefin selectivity at higher temperatures (480ºC) [25].

It was explained by Wolthuizen et al. that water has an affect on olefin selectivity

because it decreases the number of active acid sites on the catalyst as a result of competitive

adsorption between water and ethylene, which reduces the possibility of ethylene reacting further

to form other hydrocarbon species [17]. Derouane et al. made similar observations in their

investigation using 13C NMR showing that water has a higher adsorption energy than ethylene

16

and therefore can remove ethylene from the catalyst’s surface before further reactions can

commence [37]. Water could also promote desorption of light olefin species due to competitive

adsorption and possibly inhibit some bimolecular reactions such as alkylation and hydrogen

transfer [2], [38].

Deactivation and coke formation

An advantage of using ZSM-5 as a catalyst for reactions involving hydrocarbons is its

resistance to coke formation [2]. Walsh et al. investigated the difference in coke formation on

ZSM-5 and mordenite zeolites using the conversion of hexane and aromatics as model reactions

[2], [39]. They concluded that the low selectivity for coke on ZSM-5 must arise from structural

constraints on the reactions of the coke precursors. Dejaifve et al. conducted a comparative

study of the formation and stability of coke deposits using ZSM-5, offretite, and mordenite [20].

In their investigation, it was observed that carbonaceous residues are primarily formed on the

outer surface of the crystallites. This results in a slight modification of the catalyst’s molecular

shape-selective properties and also produces a high resistance to aging. It was also suggested

that the channel network of ZSM-5 contributes to the resistance to coke formation and the ease

of carbonaceous deposit removal.

Magnoux et al. investigated the deactivation modes of several catalysts through the

cracking of n-heptane (it was assumed that the cracking was the only reaction occurring) [40]. It

was observed that the initial coking of HZSM-5 is due to the “coverage” of the acid sites at

channel intersections. This initial coke is typically composed of one or two ring alkylaromatics.

As the coke content increases (above 3% coke content), polyaromatic molecules form on the

external surface and block access to a portion of the pore volume. Langner reported that the

17

amount of coke deposited on HZSM-5 depended heavily on the operating temperatures [41]. At

low temperatures in the range of 300 to 350ºC, coke precursors are formed and adsorb strongly

in the catalyst’s pore system which blocks the pores due to low mobility in the channel system.

At higher temperatures (400 to 500ºC), the mobility of the adsorbed coke precursors increases

and hydrogen transfer reactions take place. This leads to the formation of monoaromatics that

can easily desorb from the HZSM-5 catalyst surface. The evidence for these observations was

acquired by measuring increased pore volumes at higher temperatures.

The stability and catalytic action of ZSM-type zeolites as a function of Si/Al ratio was

studied by Ione et al. [42]. This was confirmed by high concentrations of aromatics at low

Si/Al ratios. As the aluminum content was decreased, the period of catalyst activity increased

and the amount of aromatic compounds decreased. Bibby et al. supported Ione et al.’s

observations by stating that the overall rate of coking is dependant on the aluminum content [43].

However, they also stated that the initial rate of coke formation and the final amount of coke

formed were not obviously linked to the aluminum content or the amount of methanol converted.

Furthermore, even after complete deactivation (essentially zero olefin production), coke

continues to accumulate on the catalyst from thermal cracking of methanol. Ducarme et al.

reached different conclusions [44]. In an investigation using ZSM-5 and ZSM-11, they

concluded that coke resistance decreased with decreasing aluminum content. The reasoning for

this observation is that deactivation due to coking occurs at the surface, which blocks entrance to

the pores, and the surface density of aluminum influences the mechanisms for coke growth on

the surface.

Bibby et al. also concluded that coke actually forms inside the catalyst channel system

with the strongest acid sites coking first, followed by weaker acid sites [43]. The presence of

18

coke inside the channels was confirmed by XRD patterns that showed structural changes in the

lattice due to the internal coke. Kärger et al observed two stages of coke formation in their

studies of n-hexane cracking [45]. During the first stage, coke is deposited in the intracrystalline

network, and in the second stage, coke is mainly deposited on the outer surface of the zeolite

crystals. Behrsing et al. analyzed HZSM-5 at different coke levels with transmission electron

microscopy (the catalyst was coked by converting methanol at 460ºC) [46]. They reported that

coke deposited inside the catalyst pores until a concentration of 6 wt% was reached. Beyond this

point, coke was deposited on the surface of the catalyst, again primarily from thermal cracking of

methanol.

Another interesting observation made by Behrsing et al. is that coke formation and

catalyst deactivation are not necessarily related when converting methanol over ZSM-5 [46].

The particular batch of ZSM-5 they were testing retained high activity (95% conversion) even

after the coke levels had reached 21 wt%. The reasoning for this is as follows: coke that was

identified to be amorphous by transmission electron microscopy was deposited in thick layers

that would still allow reactants and products to enter and exit the pores. However, the channels

become blocked and deactivation occurs when the coke contains a significant amount of

aromatic sheets that align themselves in parallel to the surface of the catalyst. Froment et al.

observed similar results in their own investigation of the cracking of mixtures of hexane with

propylene and hexadiene over ZSM-5, ZSM-11, and ZSM-48 [47]. Their results demonstrated

that the production of coke has little affect on the actual cracking reaction with ZSM-5 and ZSM-

11. As with Behrsing et al., this was due to the suggestion that the coke still allows the n-hexane

to diffuse through the pores. On the other hand, the ZSM-48, which has a one-dimensional

structure that is different than ZSM-5 and ZSM-11, suffered rapid blockage of the pores and

19

quick deactivation.

Guisnet et al. summarized the importance of coke formation on zeolites as follows [48]:

1. Coking is a shape-selective reaction. The rate of coking will be greater when (i), the

space available for the formation of coke is larger than their pore volume and (ii), the

coke precursors diffuse more slowly from the pores to the gas phase.

2. Zeolites with monodimensional pore structures or large cavities with small apertures

demonstrate a very pronounced deactivation effect from coke formation.

3. Coke deactivates zeolites by limiting or blocking access of the reactants to the active site,

not by site poisoning.

Guisnet et al. state that coking and deactivation also depend on operating conditions, reaction

time, pressure, and nature of the reactant, not solely on pore structure.

Effects of mixing states

As mentioned earlier, there has been little published work that focuses primarily on the

mixing states used for the MTO process. The literature that is available originates from

Anthony’s group at Texas A&M who compared the effect of mixing states of a continuous-

stirred tank reactor (CSTR) with a fluidized bed reactor (FBR) [49]. Anthony used 1/16”

erionite pellets at 7 atm for the CSTR and chabazite (172 µm particle size) in the FBR at 1 atm in

order to identify the appropriate means of running MTO in an industrial application. The CSTR

reactor produced very high coking rates, especially at low weight hourly space velocities of

0.2856 h-1. Due to the high rate of coking seen in MTO, Anthony concluded that the FBR

provided the most economical means of producing olefins while maintaining a constant coke

level. However, Anthony also observed that the CSTR resulted in higher ethylene selectivites

20

than was observed in the FBR. Ethylene selectivity was 56 wt% in the CSTR with a propylene

selectivity of 29 wt%. In the FBR, the propylene selectivity was about the same at 31 wt%, but

the ethylene selectivity was actually lower than for propylene (28 wt%). Additionally, C4

products were only observed in the FBR. These observations show that further investigations

into the mixing states used for the MTO process could provide additional information about the

reaction mechanisms. By simply changing the mixing states used in Anthony’s experiments, the

ethylene selectivity was changed by approximately 28 wt%. Such information could be useful

for industrial applications of MTO in terms of equipment optimization and minimization of

byproducts. While the only practical way to operate MTO is in a FBR, a FBR can be adjusted to

produce mixing states of a plug flow reactor (PFR) or a CSTR. Knowledge of which state

performs better could save industry valuable time and money by reducing the amount of

downstream processing required, plus obtaining maximum selectivity for ethylene and

propylene.

21

CHAPTER 3

EQUIPMENT AND METHODOLOGY

A “Berty” reactor and a PFR were used to provide different mixings states for this

investigation. To make a valid comparison, the following parameters had to be kept the same for

both reactors:

1. Percent conversion of oxygenates.

2. Reactor temperature.

3. Partial pressure of methanol in the feed (percent dilution).

4. Amount of coke deposited on the catalyst.

Points 1 through 3 are fairly straightforward to control. However, Point 4 presents challenges

since the small pores of the catalyst plug quickly in a PFR, resulting in rapid coke deposition. To

compensate for this difficulty, the methanol feed was diluted with helium to produce a low

partial pressure of methanol and, hence, slow the rate of coke deposition onto the catalyst.

SAPO-34 is known to provide excellent selectivity towards the light end olefins,

especially ethylene and propylene. However, it is difficult to obtain SAPO-34 for research

purposes, and synthesizing the catalyst is time consuming and produces very small yields of

catalyst. Due to these factors, ZSM-5, a known MTO catalyst, was utilized for the purpose of

examining the effect of mixing states.

For the mixing states examination, 5.0-10.0 g of ZSM-5 was first loaded into the Berty

reactor and allowed to run for a period of time. Species composition of the product gas was

determined by a gas chromatography, to measure oxygenates conversion as well as MTO product

distribution. After completion of a run, a sample of the spent catalyst from the Berty reactor was

22

analyzed by thermal gravimetric analysis (TGA) to determine coke content. Several smaller

samples (0.25 g) of the spent catalyst recovered from the Berty reactor were loaded into the PFR.

Methanol was then fed to the PFR at space velocities that resulted in the same amount of

oxygenates conversion previously obtained in the Berty reactor at the same dilution. Another

TGA was then conducted on the spent catalyst from the PFR, to make sure very little coke had

been added to the catalyst. From that, a comparison of the two mixing states from the Berty

reactor and the PFR on the MTO product distribution could be determined.

The available literature also indicates that by running at less than 100% conversion and at

low partial pressures of methanol, the selectivity for the light end olefins can be enhanced.

Higher reactor temperatures will also increase olefin selectivity. Partial pressures of methanol in

the feed ranged between 3.4 to 4.5 psia and a reactor temperature of 450ºC were used (these

conditions were used for both the PFR and Berty reactor).

Equipment

A. Apparatus

All experiments were conducted using the vaporizer/reactor/pressure regulating system

shown in Figure 3. A Hewlett-Packard model 1800A gas chromatograph detector (GCD) was

used to measure the composition of all product streams and liquid injections (See Appendix A

for calibration procedures). The GCD utilized a Supelco (2-4158) SPB™-1 sulfur column, part

#4618-02B (4.0 µm film, 30m x 0.23 mm). The temperature profile used for the analysis ran

40ºC for 3 minutes, and then ramped to 70ºC at a rate of 10ºC/min. Once 70ºC was reached, the

measurement was finished, and the column cooled back to 40ºC. Total pressure in the reactor

was controlled with a Tescom Corporation model #BB-36AL1VVA4 backpressure regulator.

23

Figure 3: Schematic of experimental setup used for the MTO mixing state investigation. The apparatus consisted of three main parts: 1-vaporizer system, 2-reactor, 3-backpressure regulating system. Points labeled A through J indicate the locations of thermocouples used to monitor the temperature of the heat tapes. All lines that made contact with the reactant feed and product stream required heating to prevent any condensation.

24

All process lines and the GCD sampling valve were wrapped with heat tapes and insulated in

order to avoid condensation of any process vapors (mainly water and methanol). Ten different

locations of the system (labeled A through J in Figure 3) were monitored with an Omega model

115KC thermocouple thermometer and type-K thermocouples.

Liquid methanol was first introduced into the vaporizer using a Cole Parmer model

74900 syringe pump. The actual vaporizer was constructed of a ¼” outer diameter stainless steel

tubing wrapped with heat tape. At the end of the vaporizer, helium and argon were mixed with

the methanol vapor, making the complete feed stream. A measured flowrate of argon was used

as an inert tracer so that the vapor flow rate exiting the reactor could be calculated using the

quantitative data from the GCD analysis. All helium and argon gases were controlled with

Matheson model 8272-0422 (200 sccm O2 calibration) mass flow controllers (MFC) and a

Matheson model 8274 multiple flow controller. The feed stream entered either the Berty reactor

or the PFR where the MTO reaction took place. At the reactor outlet, a back pressure system

was attached which maintained a constant pressure in the vaporizer and reactor, and split the

product stream so that only a very small portion of the vapor flow was fed to the GCD. The

second portion of the split (the larger of two streams) was run through a total condenser to obtain

all the condensable components from the MTO reaction. This condensate was injected into the

GCD for qualitative analysis.

As mentioned earlier, the MTO catalyst deactivates very quickly. Thus, in order to

quickly achieve steady state conditions, the vaporizer was first precharged with the

methanol/helium/argon mixture. This was achieved by closing the valve located between the

vaporizer and the reactor. The vapor stream leaving the vaporizer was vented to the fume hood

while an inert gas (helium) flowed through the reactor from a separate line (not shown). Once a

25

steady flow was obtained in the vaporizer, a series of valves was manipulated, and the feed

stream was introduced into the reactor. Approximately 1 to 5 minutes time on stream (TOS)

would pass before taking the first sample, depending on the particular experiments being

conduced. All experiments were conducted with the reactors at 450°C and 8 psig. Downstream

of the product stream split, the pressure was 2 psig (this is the pressure at the GCD sample

valve).

B. Berty Reactor

The Berty reactor was used as a CSTR for the MTO experiments (Figures 4 and 5).

Manufactured by Autoclave Engineers, the Berty reactor is a stainless steel vessel with an

internal impeller that serves as the mixer. A magnetic drive assembly located at the bottom of

the reactor was operated at 1800 RPM and drives the impeller that forces the gases to recycle

through the reactor bed, thus creating ideal back mixing. The magnetic drive is cooled with

approximately 2.0 L/min of tap water to protect the magnets from thermal damage (Figure 6). A

three-piece tube assembly was located just above the impeller inside a 5.0” diameter chamber

and serves as a catalyst basket (Figure 7). A piece of stainless steel wire mesh was placed at the

bottom of the inner most tube assembly to support the catalyst in the reactor. A stainless steel

ring compressed between two copper gaskets was used to seal the upper and lower halves of the

reaction chamber (Figure 8). The inside diameters of the upper and lower copper gaskets are

4.97” and 4.87”, respectively, and the outside diameter of both gaskets is 5.37”. All of the copper

gaskets are 0.019” in thickness. The heating jacket (Figure 9) supplied with the Berty reactor

was heated using a Lindberg temperature controller. The jacket was insulated and encased in

wire mesh to hold the insulation in place. The controller was connected to a type-K

26

Figure 4: Photograph of the Berty reactor.

27

Figure 5: Schematic of the Berty reactor. Note the flow patterns produced for the back mixing environment.

28

Figure 6: Close up of the magnetic drive system for the Berty reactor. Note the cooling lines connected to a cooling jacket located just above the magnetic assembly.

29

Figure 7: Close up of the catalyst basket unassembled (top) and assembled (bottom). This basket rests in place on the lower half of the Berty reactor just above the mixing impeller.

30

Figure 8: Copper gaskets and stainless steel spacer used to seal Berty reactor (top). Upper copper gasket and stainless steel spacer in place on the upper half of the Berty reactor (bottom).

31

Figure 9: Close up of the heating jacket used for the Berty reactor.

32

thermocouple that was attached to the heating jacket, and type-K thermocouples were placed

above and inside the catalyst basket.

C. Plug Flow Reactor

The PFR was a piece of ½” outer diameter stainless steel tubing mounted vertically in a

Lindberg 55035A furnace (Figure 10). The furnace was heated with a Barnant Company model

689-0000 temperature controller. Swagelok® fittings were used to attach the PFR to the process

lines, and quartz beads were loaded into the upper half of the reactor to aid with mixing and

heating of the feed gas. The catalyst was loaded into the bottom half of the PFR and held in

place with quartz wool. A stainless steel wire mesh tube was inserted into one end of the reactor

to hold the catalyst in place and keep the catalyst and quartz beads separated (Figure 11). The

catalyst was inside of the wire mesh tube while the quartz beads rested on top of the tube. A

type-K thermocouple was inserted into the catalyst bed through the bottom half of the PFR.

Methodology

Figure 12 illustrates a flowchart indicating the series of experiments conducted for the

MTO mixing state investigation. The ZSM-5 was first prepared via binder addition and then

tested in the PFR to check the catalyst for activity and to locate any areas in the equipment where

condensation of methanol or water could occur. The catalyst was also analyzed for coke to see

how quickly coke was deposited onto the catalyst. Next, a master batch of ZSM-5 was prepared

that would serve as the catalyst for the remainder of the investigation. After another series of

initial PFR tests, the catalyst was loaded into the Berty reactor for initial coking and analysis of

MTO product distributions in a CSTR mixing environment. The spent catalysts from the Berty

33

Figure 10: Photo of PFR secured to the upstream and downstream bulkheads. The thermocouple used to measure the temperate of the catalyst bed was inserted through the exit end of the reactor.

34

Figure 11: Stainless steel wire mesh tube that was inserted into the exit end of the reactor to hold the catalyst in place (the mesh tube is open only at one end). The catalyst was loaded inside of the wire mesh tube while the quartz beads rested on top of the mesh tube (on top of the closed end).

35

Figure 12: Flowchart of methodology for MTO mixing state investigation.

Catalyst Preparation

Initial PFR Runs

TGA Analysis

Test Equipment

Berty Runs

PFR Runs with Berty Catalyst

% Coke Product Distributions

Comparison of Mixing States

GC/MS

Catalyst Preparation

Initial PFR Runs

TGA Analysis

Test Equipment

Berty Runs

PFR Runs with Berty Catalyst

% Coke Product Distributions

Comparison of Mixing States

GC/MS

36

reactor were then analyzed for coke content and loaded into the PFR for final product

distribution measurements which were then compared to those acquired in the Berty reactor

using the same batch of catalysts.

A. Catalyst Preparation

ZSM-5 with a Si/Al ratio of 40 was used for all experiments. One kilogram of the

catalyst was obtained from Zeolyst International. The ZSM-5 arrived in a powder form and

required further preparation to make a particle size that would be compatible with the Berty

reactor (the powder form would have been carried out with the process stream otherwise).

Following the instructions provided by Zeolyst, a 30% wt SiO2 sol solution was added to the

powder to make a solution that had 20% wt silica added as a binder.

Two batches were made using the binder addition procedure. The first batch used 20 g of

ZSM-5 powder with 25 g (29.5 ml) of the sol solution. This batch was used to test the binder

addition procedure and initial PFR experiments (these initial experiments were used to test the

equipment and will be discussed later). The second batch used 500 g of ZSM-5 and 267.4 g

(736.2 ml) of the sol solution and served as the “master batch” of catalyst that was used in the

remainder of the experiments. Both batches were prepared by mixing the binder plus catalyst

until a thick, gray mixture was formed. The mixture was poured into pans with a mixture depth

of approximately 7 mm, and then placed in an oven at 105°C over night. After the catalyst was

completely dry, a brittle cake was left in the pan. This cake was chipped into large pieces, then

sieved through 3, 6, and 10 mesh sieves. The particle size obtained for the experiments was

between 3.33 and 2.00 mm.

37

B. Initial equipment/catalyst testing

Blank runs that used quartz beads as a reactor bed were made with each reactor to check

for:

1. Condensation/contamination in the reactor and the vaporizer/backpressure (V/B) system.

2. Catalytic activity of the reactor material.

3. Sufficient mixing of the gases/vapors.

After completing the blank runs, initial experiments were conducted using the first batch of

ZSM-5 catalyst. These experiments provided the following information:

1. The optimum helium dilution for running the MTO experiments.

2. A preliminary indication of what species are present in the product stream.

3. Evaluation of the binder addition procedure for preparing the ZSM-5 catalyst.

4. Evaluation of the material balance for the MTO product distribution.

To test points 1) through 4), the catalyst was loaded into the PFR as outlined in Appendix B.

C. Determination of coke levels

The amount of coke deposited onto the ZSM-5 catalyst was determined by using TGA

complimented with mass spectrometry (TGA/MS). The TGA unit (model STA 409 PC) was

manufactured by Netzsch Incorporated. The unit consists of a furnace with two power supplies

(one each for the furnace chamber and upper assembly on top of the furnace). Two rotometers

were used to control the supply of helium carrier gas and air. The air was used to burn off the

coke while the helium protected the TGA balance from vapors entering the balance chamber

from the furnace. The helium and air flow rates were 53.8 cc/min 39.2 cc/min, respectively. The

MS (model Quadstar 422) is a separate unit manufactured by Balzers Instruments. It is

38

integrated with the TGA by an insulated 0.23 mm/0.15 mm fused silica capillary column that

runs from the top of the TGA furnace to the inlet chamber of the MS. Appendix C outlines the

procedure for preparing the TGA/MS equipment for a run using the spent catalyst from the MTO

reactions. The TGA data allows for a calculation of the mass of coke deposited during the entire

MTO run. The MS measurements indicate the beginning and end of coke burn off, by measuring

the intensity of CO2 and H2O in the gas stream exiting the furnace. The MS could also indicate

if there were more that one species of coke on the catalyst (identified by changing CO2 peaks as

the temperature was ramped).

D. Base cases for the PFR and Berty reactor

A series of three runs was conducted in both the Berty reactor and the PFR. These runs

provided a standard or “base case” that could be used to periodically test the equipment and

identify any problems that may have arisen through the course of the research. The ZSM-5 from

the second batch of prepared catalyst was used for the base case runs. Conducting three runs in

each reactor also provided a measure of the repeatability of each reactor system.

E. Berty reactor experiments

The Berty reactor experiments provided the first half of the data required to make a

comparison of how mixing states affect the MTO product distributions. After completing the

initial blank runs, catalyst from the second batch of ZSM-5 (labeled “master batch”) was loaded

into the Berty reactor and used for the remainder of the experiments. The loading procedure is

outlined in Appendix B. The data collected from these twelve experiments yielded information

on the MTO product distribution in a CSTR, as a function of methanol weight hourly space

39

velocity (WHSV, defined as the mass flowrate of methanol fed per mass of catalyst or h-1),

partial pressure of methanol, TOS, and coke level. The spent catalyst also provided the loadings

for the PFR experiments that followed the Berty reactor experiments. The Berty reactor was

used for the initial coking of the ZSM-5 because it would coke the entire catalyst bed with a

constant amount of coke.

F. PFR experiments

The PFR experiments provided the other half of data necessary to make the MTO product

distribution comparisons. Using the spent catalyst from the Berty reactor, 0.25 g from each

batch was loaded into the PFR. This process was repeated three times for each batch of coked

catalyst. In order to compare the product distributions it is necessary that the following

conditions were the same in both reactor configurations:

1. Conversion of oxygenates.

2. Partial pressure of methanol in the feed stream.

3. Reactor temperature.

4. Total reactor pressure.

5. Coke levels on the catalyst.

Conditions 1) through 4) were easily adjusted. The oxygenates conversion in the PFR was

matched to the conversion in the Berty reactor by setting the methanol feed rate through trial-

and-error. The partial pressure of methanol in the feed and the total pressure of the reactor were

kept the same by adjusting the backpressure regulator and properly adjusting the helium and

argon flowrates. The reactor temperature in all experiments was controlled at 450°C.

Maintaining a constant coke level was the most difficult task with this series of experiments.

40

After following the loading and operating procedures outlined in Appendix B, each PFR

experiment was run for only one minute TOS. After one minute, a GCD sample was taken, and

the reactor was immediately flushed out with helium to quench the reaction. The catalyst was

then removed from the PFR and analyzed for coke in the TGA/MS as outlined in Appendix C.

The amount of coke measured from this last TGA/MS experiment was compared to the results

obtained from the TGA/MS runs of the spent catalyst from the Berty reactor to see if additional

coke was deposited in the PFR. If the PFR conversion was the same and there was not a

significant amount of coke added to the catalyst during the PFR run, then the one-minute TOS

GCD sample was considered to be at the same conditions as the last GCD sample taken in the

Berty reactor.

41

CHAPTER 4

RESULTS AND DISCUSSION

Initial equipment/catalyst testing

A. Equipment Testing

Blank runs were conducted in both the PFR and Berty reactor to test for:

1. Condensation/contamination in the reactor and the (V/B) system.

2. Catalytic activity of the reactor material.

3. Sufficient mixing of the gases/vapors.

A.1 Condensation/contamination in the reactor and the V/B system

Methanol and water were pumped separately through the V/B assembly at flowrates

ranging between 100 to 300 sccm without either of the reactors attached. The purpose for not

including the reactors in these first series of test was to ensure that there were no cold spots in

the process lines that would result in water/methanol condensation, which would result in

pressure “pulses”. Each test was conduced at pressures between 8 to 9 psig, and the only signs

of flow pulsing were observed at the highest water flowrates (300 sccm water). Even here, the

pulsing was small in magnitude (approximately +/- 1 psig over 1 to 2 minutes). This high water

flowrate was an extreme case that would not be used during the course of the investigation, so it

was concluded that the V/B system would perform sufficiently. Periodic samples were also

taken with the GCD to test for evidence of any reactions that may be occurring in the system.

No signs of reactions were observed at the operating vaporizer temperature of 230ºC, indicating

that there were no traces of contamination in the system. It took approximately 10 minutes for

42

the vapor stream to achieve a steady flow during the precharging phase (see Appendix B under

“Making a run with the PFR,” for clarification). The valve pressure was maintained at 2 psig to

avoid large pressure changes in the reactor when a GCD sample was taken.

A.2 Catalytic activity of the reactor material

The empty reactors were connected to the V/B system to test for catalytic activity of the

reactor materials. The V/B assembly was heated to operating temperatures and connected to the

PFR, which was first heated to 250ºC. A mixture consisting of 83.3% argon and 16.7%

methanol (250 sccm argon, 50 sccm methanol) was fed through the entire reactor system. This

particular mixture was chosen because it was the starting point for deciding what helium

dilutions should be used for the MTO experiments. Samples were taken every 11 minutes to

check for indications of reaction with the methanol. It was only at 450ºC that traces of

formaldehyde were visible and assumed negligible. However, this breakdown of methanol was

enough to leave a light film of coke on the reactor walls and on the quartz beads, but was easily

removed with deionized water and a test tube brush. The quartz beads were cleaned by heating

the beads in an oven to 600ºC and holding for two hours to burn off the coke. This same reactor

activity test was used with the Berty reactor, and produced similar results; trace quantities of

formaldehyde, only at an operating temperature of 450ºC.

A.3 Gas/vapor mixing tests

To address the mixing of gases/vapors in the PFR, three different argon/methanol

mixtures were injected into the reactor system and analyzed with the GCD. This was to ensure

that the feed gases were properly mixed before entering the reactor. The data in Table 1 presents

43

Table 1: Flow parameters and results for PFR mixing tests (blank, no catalyst)

Ar/MeOH Gas Flows [sccm] Average measured mol fraction Test # Composition Ar MeOH Ar MeOH

1 50/50 100 100 0.487 (+/- 2.48%) 0.513 (+/- 2.36%) 2 77/23 100 332 0.748 (+/- 1.61%) 0.252 (+/- 4.77%) 3 50/50 200 200 0.477 (+/- 1.06%) 0.523 (+/- 0.97%)

44

the results obtained from the GCD during these mixing tests in the PFR. The largest relative

error calculated was less than 5.0%, indicating that there was minimal deviation and sufficient

mixing in the system (each of the averages were calculated with 6 to 10 values). The data also

indicated that the calculated methanol concentration was slightly higher than fed to the reactor.

This was probably due to the presence of tails on the methanol peaks measured with the GCD.

The degree of mixing for the Berty reactor was examined while conducting the base case

experiments and will be discussed in the section titled “ZSM-5 base cases”.

B. Catalyst testing

The purpose of the next series of experiments was to acquire the following information

about the reaction system:

1. Optimum helium dilution for running the MTO experiments.

2. The species that are present in the product stream.

3. Evaluation of the binder addition procedure for preparing the ZSM-5 catalyst.

4. The degree to which the material balance could be closed.

The parameters and data analysis for all runs conducted with ZSM-5 (both initial and

master batches) are located in Appendix D. Helium dilutions of 90%, 85%, and 80% in the

oxygenate feed were tested in order to establish the highest dilution that can be used and still

acquire measurable samples (sufficient signal to noise ratios). These three dilutions were

conducted with a methanol flowrate of 5.0 WHSV (29.1 sccm), a reactor pressure of 8 psig, and

a catalyst bed temperature of 450ºC. While olefin products were observed for each of these

dilutions, at 90% dilution the signal to noise ratio was so low that the C4= product concentration

could not be quantified. Both 85% and 80% helium dilution produced sufficient signal strength,

45

but the 85% dilution was chosen to provide a slower rate of coking. The results from these tests

also proved that the ZSM-5 catalyst was active in the conversion of methanol to olefins, hence

the binder addition procedure described in Chapter 3 was successful.

Next, a series of experiments were conducted to determine the effect of using different

space velocities for the MTO reaction. Table 2 lists the parameters that were used and the results

are presented in Figures 13-17. Ethylene, butene, and saturates (only C4 saturates) are essentially

unaffected by space velocity with measured mole fractions of approximately 0.28 ± 0.02, 0.10 ±

0.01, and 0.02 ± 0.00, respectively. There were no traces of methane or carbon dioxide.

However, variations in the mole fractions are observed with propylene and carbon monoxide.

Examination of Figure 14 indicates that the mole fraction of propylene increased from 0.30 to

0.47 ± 0.04 as the methanol space velocity increased from 5.0 to 15.05 WHSV. This is in

agreement with the literature that states that a decrease in methanol conversion (produced by

increasing WHSV) increases olefin selectivity, in this case, propylene selectivity [21], [22]. At

the lower space velocities, more propylene was allowed to form aromatic products, and hence,

lower propylene concentrations. The higher coke levels at low space velocities, as shown in

Table 2, would also indicate the aromatization of propylene.

Higher space velocities, or shorter space times, reduces the degree to which the

series/parallel reaction system continues to produce aromatic species. Examining the carbon

monoxide data from Figure 17, which shows a decrease in carbon monoxide with increased

methanol WHSV, reinforces this argument. As the methanol flow increased from 5.0 to 15.0

WHSV, carbon monoxide formed by the following equation was reduced:

COH2OHCH 23 +→ (4)

46

Table 2: Flow parameters used for examining the effects of space velocity on the MTO product distribution. Reactor at 450ºC, 8 psig, sample valve at 2 psig, 85% dilution, 0.50 g ZSM-5 (initial batch).

Run WHSV TOS Total Final % CokeNumber [h-1] [min] Helium Methanol [sccm] [wt %]

6 5.00 113 164.8 29.2 194.0 2.058 9.03 103 297.3 52.6 349.9 1.58

10 15.05 105 495.5 87.7 583.2 1.27

Flowrates [sccm]

47

Figure 13: Mole fraction of ethylene versus time on stream for 5.00, 9.03, and 15.05 WHSV using 0.50 g ZSM-5 (initial batch). Reactor at 450ºC and 8 psig, sample valve at 2 psig, 85% helium dilution. [Water, dimethyl ether, methanol free basis]

0.00

0.10

0.20

0.30

0.40

0.50

0.60

0.70

0.80

0.90

1.00

0 20 40 60 80 100 120Time on Stream (min)

5.00 WHSV9.03 WHSV15.05 WHSV

Mol

Fra

ctio

n of

Eth

ylen

e

48

Figure 14: Mole fraction of propylene versus time on stream for 5.00, 9.03, and 15.05 WHSV using 0.50 g ZSM-5 (initial batch). Reactor at 450ºC and 8 psig, sample valve at 2 psig, 85% helium dilution. [Water, dimethyl ether, methanol free basis]

0.00

0.10

0.20

0.30

0.40

0.50

0.60

0.70

0.80

0.90

1.00

0 20 40 60 80 100 120Time on Stream (min)

5.00 WHSV9.03 WHSV15.05 WHSV

Mol

Fra

ctio

n of

Pro

pyle

ne

49

Figure 15: Mole fraction of butene versus time on stream for 5.00, 9.03, and 15.05 WHSV using 0.50 g ZSM-5 (initial batch). Reactor at 450ºC and 8 psig, sample valve at 2 psig, 85% helium dilution. [Water, dimethyl ether, methanol free basis]

0.00

0.10

0.20

0.30

0.40

0.50

0.60

0.70

0.80

0.90

1.00

0 20 40 60 80 100 120Time on Stream (min)

5.00 WHSV9.03 WHSV15.05 WHSV

Mol

Fra

ctio

n of

But

ene

50

Figure 16: Mole fraction of saturates versus time on stream for 5.00, 9.03, and 15.05 WHSV using 0.50 g ZSM-5 (initial batch). Reactor at 450ºC and 8 psig, sample valve at 2 psig, 85% helium dilution. [Water, dimethyl ether, methanol free basis]

0.00

0.10

0.20

0.30

0.40

0.50

0.60

0.70

0.80

0.90

1.00

0 20 40 60 80 100 120Time on Stream (min)

5.00 WHSV9.03 WHSV15.05 WHSV

Mol

Fra

ctio

n of

Sat

urat

es

51

Figure 17: Mole fraction of carbon monoxide versus time on stream for 5.00, 9.03, and 15.05 WHSV using 0.50 g ZSM-5 (initial batch). Reactor at 450ºC and 8 psig, sample valve at 2 psig, 85% helium dilution. [Water, dimethyl ether, methanol free basis]

0.00

0.10

0.20

0.30

0.40

0.50

0.60

0.70

0.80

0.90

1.00

0 20 40 60 80 100 120Time on Stream (min)

5.00 WHSV9.03 WHSV15.05 WHSV

Mol

Fra

ctio

n of

CO

52

The actual methanol conversion could not be measured at this point in the investigation

because the methanol peak required recalibration to compensate for peak tailing (this is

addressed later when discussing the material balance). However, Figures 18-19 show an

interesting trend in terms of what happens to the oxygenates. As time progresses, DME reaches

a constant mole fraction of 0.10 ± 0.01, leaving methanol the only oxygenate that increases in

concentration as the catalyst deactivates. Since the olefin distribution remains constant over the

course of a run, this implies that there are two possible reaction paths for methanol to form

olefins. The first is the standard reaction path presented in Equation 1 where methanol first

forms DME before producing the olefins. The second reaction path could be the direct

conversion of methanol to olefin species.

The next step was to acquire qualitative information pertaining to the higher carbon

species in the product stream. Liquid condensate collected from Runs 8 and 10 were analyzed

with the GCD by injecting 1.0 µL of each sample into the column via the GC septum. The

column was maintained at 60ºC for three minutes, and then ramped to 200ºC at 5ºC/min. The

peaks obtained were extracted by the GCD software and identified as aromatic compounds in the

C8-C10 range (species such as ethylbenzene, toluene, xylene). This was similar to the types of

heavy carbon species reported in previous literature when running MTO with ZSM-5 [3], [4],

[13]. Ideally, the operating conditions chosen for this investigation should optimize the olefin

selectivites such that the amount of aromatic species is reduced, and the data may be normalized

without the aromatic species. A carbon balance was used to determine the significance of the

larger aromatic compounds and decide if normalizing the data will work for this investigation.

However, up until this point it was observed that the methanol peaks measured with the GCD

were much smaller than originally calibrated for. Large tails were visible in the scans and this

53

Figure 18: Mole fraction of dimethyl ether versus time on stream for 5.00, 9.03, and 15.05 WHSV using 0.50 g ZSM-5 (initial batch). Reactor at 450ºC and 8 psig, sample valve at 2 psig, 85% helium dilution. [Water free basis]

0.00

0.10

0.20

0.30

0.40

0.50

0.60

0.70

0.80

0.90

1.00

0 20 40 60 80 100 120Time on Stream (min)

5.00 WHSV9.03 WHSV15.05 WHSV

Mol

Fra

ctio

n of

DM

E

54

Figure 19: Mole fraction of methanol ether versus time on stream for 5.00, 9.03, and 15.05 WHSV using 0.50 g ZSM-5 (initial batch). Reactor at 450ºC and 8 psig, sample valve at 2 psig, 85% helium dilution. [Water free basis]

0.00

0.10

0.20

0.30

0.40

0.50

0.60

0.70

0.80

0.90

1.00

0 20 40 60 80 100 120Time on Stream (min)

5.00 WHSV9.03 WHSV15.05 WHSV

Mol

Fra

ctio

n of

MeO

H

55

made it difficult to determine the actual concentration of methanol in the system.

In order to correct for the tailing observed while diluting the gas stream with helium, the

calibration constants for methanol were determined for a range of methanol areas measured with

the GCD. Table 3 lists the experimental conditions for these experiments, along with the final

constants that were used for the remainder of the investigation. The constants were determined

as follows: known concentrations of argon and methanol were injected into the PFR reactor,

which was heated to 250ºC (this ensures that there are no cold spots and no thermal breakdown

of methanol into formaldehyde). With the calibration constant for argon already determined

during preliminary GCD calibrations, the solver function was used in Microsoft® Excel 2000 to

back out what the required calibration constant for methanol should be to calculate the preset

concentrations of each component. The constants listed in Table 3 were programmed into the

data analysis spreadsheet using Boolean operators (the Microsoft® Excel 2000 Boolean code can

be found in Appendix A). Care was taken for the remainder of the investigation to check that the

measured methanol areas fell within the area ranges used to determine the calibration constants.

To determine the degree to which the material balance could be closed, 0.5 g of ZSM-5 was

loaded into the PFR using a methanol flow of 52.6 sccm at 85% dilution to give 9.03 WHSV.

The gas mixture used to achieve the 85% dilution was 9.8 sccm argon and 288.7 sccm helium,

resulting in a feed composition of 2.8% argon, 82.2% helium, and 15.0% methanol. The argon

was used as a tracer and provided a reference necessary for calculating the total flow of products

plus unreacted oxygenates (minus water). The reactor was operated at 450ºC and 8 psig with the

sample valve at 2 psig and the experiment was run for 104 minutes TOS. The data acquired

from this run and all calculations for the material balance can be found in Appendix E.

Typically, during the MTO reaction there is an “initial coking” of the strongest acid sites before

56

Table 3: Flow parameters and acquired methanol calibration constants for recalibration of methanol at 85% helium dilution (Blank run, no catalyst).

Calibration Total Composition Constant DeterminedTest Helium Ar Methanol [sccm] Ar/MeOH For Methanol1C 300.1 27.3 26.8 354.2 .505/.495 26633542C 300.1 44.1 5.0 350.2 .898/.102 5810743C 303.1 37.9 16.2 357.2 .701/.299 25452274C 301.1 42.2 10.8 354.1 .796/.204 1779574

Flowrates [sccm]

57

a steady rate of coking is achieved. For this reason, data from the first sample was not

considered for the material balance. Instead, data taken after 16 minutes TOS was chosen since

it would be at a period of high catalytic activity (this is important to avoid false indications of

conversions from thermal breakdown of methanol to formaldehyde after the catalyst is

thoroughly coked) and is probably the closest to being at a steady state coke level in the PFR

experiments. Carbon lost due to coking of the catalyst and the reactor walls was not included in

the balance.

Completion of the material balance calculations indicated that out of 2.351 x 10-3

mol/min of carbon fed to the reactor, approximately 1.797 x 10-3 mol/min of carbon was reacted

producing 1.053 x 10-3 mol/min of products. This leaves 5.540 x 10-4 mol/min, or 23.7% of the

carbon unaccounted for. Most of this unaccounted carbon is ignored because only the lighter

hydrocarbon products (these appear in the GCD scans after 5 minutes into the measurement) are

analyzed for concentration. The larger carbon species (C8-C10 compounds) are neglected in the

normalization of the product distribution, but can be seen in the GCD scans if the measurement

is run for longer time durations. Assuming that the C9 carbon species is the most abundant heavy

hydrocarbon (use these 9 carbons as the basis for the calculation of the remaining carbon

species), the molar flowrate of C9 species was calculated and compared to the total molar

flowrate of products that are actually measured with the GCD. With this assumption,

approximately 6.156 x 10-5 mol/min of the higher carbon products, or roughly 5.5% of the total

carbon species are the C8-C10 compounds unaccounted for. These results, being less than 10% of

the total molar flowrate of products, were considered to be satisfactory in completing the

material balance.

58

Base cases for the PFR and Berty reactor

Three separate catalyst loadings, all from the same “master batch” described in Chapter

3, were used to obtain three replicate runs for both the PFR and Berty reactor. The PFR used

0.50 g of ZSM-5 per run while the Berty reactor was loaded with 10.00 g of ZSM-5 per loading.

Table 4 lists the experimental parameters used for each of the 6 base case runs. All runs were

conducted at 450ºC and 8 psig (reactor conditions) with the sample valve at 2 psig.

Figures 20-22 present the results from the PFR base case runs. The mole fractions for

ethylene, propylene, butene, and saturates were approximately 0.25 ± 0.02, 0.42 ± 0.04, 0.10 ±

0.01, and 0.02 ± 0.00, respectively. Carbon monoxide and carbon dioxide mole fractions were

0.19 ± 0.06 and 0.01 ± 0.01, respectively. These six mole fractions were all calculated on a

water, argon, DME, and methanol free basis and were found to be similar to the results obtained

during the initial PFR runs using 9.03 WHSV (shown earlier in Figures 13-19). Table 5

compares the data collected from the initial and base case PFR runs and indicates that the

product distribution between the two series of experiments were similar. The only noticeable

difference was a trace of carbon dioxide that was most likely produced from the following water-

gas shift reaction:

222 HCOOHCO +⇔+ (5)

On a water and argon free basis, the mole fractions of methanol and DME in the products were

approximately 0.37 to 0.48 ± 0.05 and 0.15 ± 0.01, respectively. There were no measurable

amounts of methane, and the saturates consisted of only iso-butane and n-butane. Examination

of Figure 20 indicates that the mole fractions of olefin products exiting the reactor were constant

59

60

Figure 20: Average mole fraction of light olefins and saturates versus time on stream for the three PFR base case runs. Reactor at 450ºC and 8 psig, sample valve at 2 psig, 85% total dilution, 9.03 WHSV, 0.50 g ZSM-5 (master batch). [Water, argon, dimethyl ether, methanol free basis]

0.00

0.10

0.20

0.30

0.40

0.50

0.60

0.70

0.80

0.90

1.00

0 10 20 30 40 50 60 70Time on Stream (min)

Ethene

Propene

Butenes

SaturatesMol

Fra

ctio

n

61

Figure 21: Average mole fraction of carbon monoxide and carbon dioxide versus time on stream for the three PFR base case runs. Reactor at 450ºC and 8 psig, sample valve at 2 psig, 85% total dilution, 9.03 WHSV, 0.50 g ZSM-5 (master batch). [Water, argon, dimethyl ether, methanol free basis]

0.00

0.10

0.20

0.30

0.40

0.50

0.60

0.70

0.80

0.90

1.00

0 10 20 30 40 50 60 70Time on Stream (min)

CO

CO2

Mol

Fra

ctio

n

62

Figure 22: Average mole fraction of methanol and dimethyl ether with total oxygenates conversion versus time on stream for the three PFR base case runs (averaged). Reactor at 450ºC and 8 psig, sample valve at 2 psig, 85% total dilution, 9.03 WHSV, 0.50 g ZSM-5 (master batch). [Water, argon free basis]

0.00

0.10

0.20

0.30

0.40

0.50

0.60

0.70

0.80

0.90

1.00

0 20 40 60 80TOS (min)

0%

10%

20%

30%

40%

50%

60%

70%

80%

90%

MeOH

DME

OxygenateconversionM

ol F

ract

ion

% C

onversion

63

Table 5: Product comparison between the initial (Run 8) and base case PFR runs for 9.03 WHSV. Reactors at 450ºC, 8 psig, sample valve at 2 psig, 85% dilution, 0.50 g ZSM-5 used in each data series. [Water, argon, dimethyl ether, methanol free basis]

Product Initial PFR Base Case Error (+/-)Ethylene 0.28 0.25 0.02

Propylene 0.38 0.42 0.04Butene 0.10 0.10 0.01

Saturates 0.02 0.02 0.00CO 0.25 0.19 0.06CO2 0.00 0.01 0.01

TOS [min] 103 60 -wt% coke 1.58 1.34 -

Mol Fraction of products

64

over the course of 60 minutes TOS. This was reasonable due to the high dilution (85%) of the

methanol feed and the short TOS that resulted in a slow rate of deactivation as seen in Figure 22.

The average oxygenates conversion only dropped from 77.1% to 65.8% conversion over 60

minutes. It was again observed that the concentration of DME remained constant as the methanol

concentration increased during the course of the run. This again supports the theory that there

are two possible reaction paths for olefin production.

Figures 23-25 present the results from the Berty reactor base case runs. The average

mole fractions for ethylene, propylene, butene, and saturates were approximately 0.28 ± 0.02,

0.39 ± 0.01, 0.10 ± 0.00, and 0.23 ± 0.02, respectively. The average carbon dioxide mole

fraction was 0.02 ± 0.00 formed by the water-gas shift reaction (Equation 5) discussed earlier.

These five mole fractions were all calculated on a water, argon, DME, and methanol free basis.

On a water and argon free basis, the mole fractions of methanol and DME in the products were

approximately 0.40 ± 0.04 and 0.02 ± 0.00, respectively. There were no measurable amounts of

methane or carbon monoxide, and saturates consisted of propane, iso-butane and n-butane. A

constant product distribution was observed over 291 minutes TOS in addition to a constant

average oxygenates conversion of 84.4%. Table 6 lists the product distribution data obtained

from the PFR and Berty base cases plus measured coke levels and total oxygenates conversion.

The amount of saturates was 11.5 times larger in the Berty reactor than observed in the PFR. In

addition, approximately 80% of saturates in the Berty reactor was propane, which could be due

to a low concentration of oxygenates (the concentration of oxygenates in the Berty reactor is

much lower that in the PFR because the reactants immediately go to final concentrations, which

are very low in the mixing environment that is diluted with helium and water). Propane was

never observed during the PFR base case runs and may be another result of the low space

65

Figure 23: Average mole fraction of light olefins and saturates versus time on stream for the three Berty reactor base case runs. Reactor at 450ºC and 8 psig, sample valve at 2 psig, 80% total dilution, 0.903 WHSV, 10.0 g ZSM-5 (master batch). [Water, argon, dimethyl ether, methanol free basis]

0.00

0.10

0.20

0.30

0.40

0.50

0.60

0.70

0.80

0.90

1.00

0 50 100 150 200 250 300Time on Stream (min)

Ethene

Propene

Butenes

SaturatesMol

Fra

ctio

n

66

Figure 24: Average mole fraction of carbon dioxide versus time on stream for the three Berty reactor base case runs. Reactor at 450ºC and 8 psig, sample valve at 2 psig, 80% total dilution, 0.903 WHSV, 10.0 g ZSM-5 (master batch). [Water, argon, dimethyl ether, methanol free basis]

0.00

0.10

0.20

0.30

0.40

0.50

0.60

0.70

0.80

0.90

1.00

0 50 100 150 200 250 300Time on Stream (min)

CO2

Mol

Fra

ctio

n

67

Figure 25: Average mole fraction of methanol and dimethyl ether with total oxygenates conversion versus time on stream for the three Berty reactor base case runs. Reactor at 450ºC and 8 psig, sample valve at 2 psig, 80% total dilution, 0.903 WHSV, 10.0 g ZSM-5 (master batch). [Water, argon free basis]

0.000.100.200.300.400.500.600.700.800.901.00

0 50 100 150 200 250 300

TOS (min)

0%10%20%30%40%50%60%70%80%90%100%

MeOH

DME

Oxygenateconversion

Mol

Fra

ctio

n % C

onversion

68

Table 6: Product comparison between the PFR and Berty base case runs. Reactors at 450ºC, 8 psig, sample valve at 2 psig. PFR loaded with 0.50 g ZSM-5 (master batch) with 85% feed dilution. Berty reactor loaded with 10.0 g ZSM-5 (master batch) with 80% feed dilution. [Water, argon, dimethyl ether, methanol free basis except for dimethyl ether and % oxygenates conversion (these are on a water, argon free basis)]

Product PFR Berty Error (+/-)Ethylene 0.25 0.28 0.02

Propylene 0.42 0.39 0.04Butene 0.10 0.10 0.01

Saturates 0.02 0.23 0.00CO 0.19 0.00 0.06CO2 0.01 0.02 0.01DME 0.15 0.02 0.01

% OxygenatesConversionTOS [min] 60 291 -wt% Coke 1.34 0.88 -

Mol Fraction of products from base case studies

77.1 3.684.4

69

velocities used. Another significant difference in the Berty product distribution was that there

was never any carbon monoxide (there was always carbon monoxide present in the initial and

base case PFR runs). An initial stage of coking may occur, similar to what was observed with

the PFR. This would again be attributed to coking the strongest acid sites first.

It should be brought to attention that the dilution used for the Berty reactor for the base

case studies was 80% as opposed to the 85% dilution used in the PFR. The dilution was lowered

in the Berty reactor due to the difficultly in achieving the high helium flowrates required for the

larger catalyst loadings that were used in these experiments. However, the remainder of the

experiments related to the product distribution comparisons were all conducted at 80% total

dilution, both in the Berty reactor and the PFR.

Berty reactor experiments

Four “master” batches of coked catalysts were produced in the Berty reactor. Table 7

summarizes the experimental parameters used to produce the batches of coked catalyst. At the

start of each run, the catalyst was subjected to the same conditions used for the Berty base case

experiments. This ensured that everything was in working order before proceeding on with the

scheduled experiments.

A difficulty encountered with this series of experiments was the attempt to achieve high

levels of coke on the ZSM-5 catalyst. As indicated in the literature, ZSM-5 is resistant to the

formation of coke [18], [19], [20]. In addition, the catalyst is even less prone to coking in the

Berty reactor because of the CSTR mixing conditions. According to CSTR theory, the

concentration of reactants and products immediately reach their final concentrations and the first

four runs conducted in the Berty reactor were operated at conversions of 70% and higher. This

70

Table 7: Experimental parameters and measured coke content of Berty runs that were used in comparison PFR experiments.

Run Wt % TOS Oxygenates WHSV Mass of MeOH FlowNumber Coke [min] Conversion [h-1] catalyst [g] [g/hr]

7 1.39 506 81.8% 0.903 10.0 9.039 2.40 556 38.2% 4.000 5.0 20.0010 2.60 535 69.6% 0.700 5.0 3.5011 1.41 114 36.1% 8.000 5.0 40.00

71

translates into very low methanol and DME concentrations in the system and a very slow rate of

coking. To compensate for this problem, the helium flow was shut off after running the initial

check against the Berty base case runs. The methanol flowrate was then set for 1.0 WHSV

(105.3 sccm) and GCD samples were taken once an hour to monitor the reaction and look for

signs of catalyst deactivation. This procedure of coking allowed for a sufficient level of coke to

be deposited on the catalyst and to ensure that a sufficient fraction of the catalyst’s active sites

were blocked so as to produce variances in product distribution. After 4-8 hours, the helium

flow was re-established and the feed mixture was adjusted to achieve a targeted amount of

conversion, at which point, the final product distribution was measured. Three samples were

taken at the end of each run to insure steady state reaction conditions. Using this procedure, four

master batches of coked catalyst were obtained with their coke levels and CSTR conversions

listed in Table 7.

Approximately 120 to 130 mg of the spent catalyst from the runs in Table 7 were

analyzed for coke content using the TGA/MS, as discussed in Chapter 2 and Appendix C.

Figure 26 illustrates an example of a completed TGA/MS scan. The coke on the catalyst began

to burn off at about 350ºC and the completion of burn-off was indicated by the leveling of the

mass and the CO2 traces. An interesting point is that two major CO2 peaks are visible during the

TGA analysis. This indicates that there was probably more than one type of coke species on the

catalysts. This is in agreement with literature in which both paraffinic and aromatic coke

deposited on the catalyst during MTO have been reported [43].

Figures 27-28 depict the final product distribution for the four batches of coked catalyst

from the Berty reactor. The products consisted of light olefins and saturates with no measurable

traces of methane, carbon monoxide, or carbon dioxide. As the methanol WHSV was increased,

72

Figu

re 2

6: E

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of a

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alys

is (F

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Run

10)

. Th

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furn

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200

ºC fo

rre

mov

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ater

, the

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ld a

t 800

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The

plot

con

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s bot

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A d

ata

plus

the

MS

data

for

carb

on d

ioxi

de. T

he C

O2 p

eak

was

use

d to

indi

cate

the

begi

nnin

g an

d en

d of

cok

e re

mov

al.

The

first

mas

s cha

nge

is w

ater

rem

oval

; the

seco

nd m

ass c

hang

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(Fee

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% O

2)

73

Figure 27: Mole fraction of light olefins and saturates for the four coked master batches of catalyst from the Berty reactor. Reactor at 450ºC and 8 psig, sample valve at 2 psig, 80% total dilution. Coke levels are 1.39, 1.41, 2.60, and 2.40 wt% for 81.8, 36.1, 69.6, and 38.2% oxygenate conversion, respectively. [Water, argon, dimethyl ether, methanol free basis]

0.000.050.100.150.200.250.300.350.400.450.500.55

Ethylene

Propyle

ne

Butene

Satura

tes

81.8% Conversion36.1% Conversion69.6% Conversion38.2% Conversion

Mol

Fra

ctio

n

74

Figure 28: Mole fraction of methanol and dimethyl ether for the four coked master batches of catalyst from the Berty reactor. Reactor at 450ºC and 8 psig, sample valve at 2 psig, 80% total dilution. Coke levels are 1.39, 1.41, 2.60, and 2.40 wt% for 81.8, 36.1, 69.6, and 38.2% oxygenate conversion, respectively. [Water, argon free basis]

0.00

0.10

0.20

0.30

0.40

0.50

0.60

0.70

0.80

0.90

DME

MeO

H

81.8% Conversion36.1% Conversion69.6% Conversion38.2% Conversion

Mol

Fra

ctio

n

75

there was a small decrease in ethylene mole fraction (0.33 to 0.27) and an increase in butene

mole fraction (0.11 to 0.15), for the low coked catalyst. The most significant changes, however,

were in propylene and saturate concentrations. The mole fraction of propylene increased from

0.37 to 0.52 and the saturates decreased from 0.19 to 0.06. The magnitude and direction of

these changes implies that propylene is a main precursor for saturate production in the Berty

reactor, which is suppressed by increasing the methanol WHSV. As can be seen in Figure 28,

the oxygenates concentration was equivalent to methanol conversion, although the DME

increased with both coke level and with decreasing oxygenate conversion.

As conversion was decreased for the high coke catalyst (2.60 and 2.40 wt% coke), there

was an increase in ethylene (0.24 to 0.36) and a decrease in butene (0.14 to 0.07). This was

opposite of that observed with the low coke catalyst in the Berty reactor. The higher coke levels

may have altered the reaction path of the olefins to result in these changes in ethylene and butene

mol fractions. There was still an increase in propylene mole fraction (0.47 to 0.52) accompanied

by a decrease in saturates (0.17 to 0.05), which indicates that propylene selectivity is only

dependent on coke levels, at high conversions. The behavior of propylene and saturates

observed for the low and high coked catalysts would indicate that the hydrocarbon pool

mechanism shown in Figure 1 could be better represented by the mechanism shown in Figure 29

where the majority of saturates and coke originate from propylene. This would then account for

the relationship between propylene and propane observed in the Berty reactor.

Examination of the Berty runs that were exposed to high methanol partial pressures for

coking purposes (Runs 20 through 24 in Appendix D) indicates several interesting trends. As the

partial pressure of methanol was increased from 4.5 to almost 22.2 psia, a substantial increase in

saturates concentration at 1.00 WHSV was observed (the saturates were over 50% of the

76

Figure 29: Schematic of the corrected hydrocarbon pool mechanism.

77

measured hydrocarbon products). This was to be expected, as it is reported in the literature that

high partial pressures of methanol reduces overall olefin selectivity [21], [25], [32], [33].

Approximately 65% of the saturates was propane, similar to that observed for the runs listed in

Table 7. Another observation was that increasing WHSV’s (2.0 to 8.0 h-1) resulted in decreased

total oxygenate conversions and reduced saturates in the product stream (approximately 30% of

total products). The fraction of propane in the saturates was reduced to approximately 50% of

the total saturates, which was seen also in the runs from Table 7.

Carbon monoxide was never observed in the Berty reactor, regardless of coke level,

WHSV, or methanol partial pressure. This could be attributed solely to the characteristic of the

CSTR mixing environment.

PFR experiments

The four batches of coked catalyst obtained in the Berty reactor were used to

complete the PFR experiments for the final stage of this investigation. Table 8 tabulates the PFR

experiments conducted at high conversions. The coked catalysts from the Berty runs were

exposed to a methanol feed stream containing 80% helium/argon for dilution, and the PFR was

operated at 450ºC and 8 psig, with the sample valve at 2 psig. All of the PFR results shown in

Table 8 correspond to an exit gas sample taken at TOS = 1 minute, to insure that the coke levels

would be nearly identical to the levels obtained in the Berty. Runs 29-31 resulted in an average

conversion of 70.0 ± 1.2%, demonstrating satisfactory repeatability for the short TOS. Coke

analyses of the spent PFR catalyst (i.e., after TOS = 1 minute) showed only a 1-5% increase in

the coke level. As can be seen, Run 27 produced a nearly identical conversion with that

measured in Berty reactor. Runs 33-35 were averaged in the case of the 2.60 wt% coked

78

Table 8: Experimental parameters and measured conversions for PFR runs that used coked catalyst from the Berty reactor. Final conversions measured at 80% dilution with 10.0 sccm argon in mixture. All runs used 0.25 g of catalyst per loading. Time on stream for each run before sampling was 1 minute.

Original New Run Final PFR WHSV MeOH FlowRun Number Number Conversion [h-1] [g/hr]

[21] 26 89.0% 9.03 2.26 1.39 wt% coke 27 81.4% 11.00 2.75

81.80% 28 87.5% 11.00 2.75Conversion 29 69.9% 15.00 3.75

30 71.7% 15.00 3.7531 69.3% 15.00 3.75

[24] 32 81.3% 5.00 1.252.60 wt% coke 33 61.2% 10.00 2.50

69.60% 34 60.0% 10.00 2.50Conversion 35 65.0% 10.00 2.50

79

catalysts to give an average oxygenate conversion of 62.1 ± 2.6%, and the average of these runs

was used for the PFR/Berty comparison.

Before comparing the effect of mixing states on the MTO product selectivites, it is

interesting to review the results obtained with all of the PFR runs which employed spent

catalysts from the Berty rector. The complete data set for these experiments can be found in

Appendix D. Examination of Figures 30 and 31 indicates that the ratio of ethylene to propylene

becomes greater than 1.0 somewhere below 80% conversion. That is, as the conversion of

oxygenates decreases, ethylene becomes the most abundant olefin in the product distribution. As

with all the PFR runs, the DME concentrations were essentially constant, independent of the

oxygenate conversions. The data were compared in Figure 32 and show that as the conversion

was reduced for both the low and high coked Berty catalyst and the ethylene mol fraction

increased significantly while propylene decreased to where the ethylene to propylene ratio was

greater than 1.0. The ratio was always less than 1.0 for PFR studies using fresh ZSM-5. This

could indicate a change in the reaction environment on the catalyst caused by previous use in the

Berty reactor.

Another unexpected occurrence in these experiments was the increased production of

carbon dioxide in all of the PFR runs listed in Table 8. Again, this would indicate that some

physical characteristic of the catalyst changed after use in the Berty reactor. It could very well

be an effect of the mixing environment on coke deposition. This, in combination with the

absence of carbon monoxide in the product gas, suggests that the water gas shift reaction occurs

at a high efficiency over these coked catalysts. One possibility is that the coke precursors in the

Berty rectors are different than in the PFR, leading to an altered adsorption pattern that is

beneficial to the water gas shift reaction. The type of coke that was deposited in the Berty

80

Figure 30: Mole fraction of light olefins and saturates for PFR runs using the 1.39 wt% coked catalyst from the Berty reactor. Reactor at 450ºC and 8 psig, sample valve at 2 psig, 80% total dilution, 0.25 g ZSM-5. The data points at 70% oxygenates conversion are averages of PFR runs 3-5. [Water, argon, dimethyl ether, methanol free basis]

0.000.050.100.150.200.250.300.350.400.450.500.55

CO2

Ethylene

Propyle

ne

Butene

Satura

tes

89.0% Conversion81.4% Conversion70.0% Conversion

Mol

Fra

ctio

n

81

Figure 31: Mole fraction of light olefins and saturates for PFR runs using the 2.60 wt% coked catalyst from the Berty reactor. Reactor at 450ºC and 8 psig, sample valve at 2 psig, 80% total dilution, 0.25 g ZSM-5. The data points at 62.1% oxygenates conversion are averages of PFR runs 2-4. [Water, argon, dimethyl ether, methanol free basis]

0.000.050.100.150.200.250.300.350.400.450.500.55

CO2

Ethylene

Propyle

ne

Butene

Satura

tes

81.3% Conversion62.1% Conversion

Mol

Fra

ctio

n

82

Figure 32: Ethylene and propylene mole fraction comparison between the coked Berty catalyst and fresh catalyst (Fresh catalyst data is from Run 8). Conversions for the 1.39 and 2.60 wt% coke are 81.4 and 81.3% (above) and 70.0 and 62.1% (below). Conversion was not measured for Run 8. [Water, argon, dimethyl ether, methanol free basis]

0.000.050.100.150.200.250.300.350.400.450.500.55

Fresh Catalyst 1.39 wt% Coke 2.60 wt% Coke

EthylenePropylene

Mol

Fra

ctio

n

0.000.050.100.150.200.250.300.350.400.450.500.55

Fresh Catalyst 1.39 wt% Coke 2.60 wt% Coke

EthylenePropylene

Mol

Fra

ctio

n

83

reactor may be different in composition than produced in the PFR, or it could cover or attach to

the active sites in the catalyst differently.

Comparison of product distributions

The data collected from the PFR and Berty reactor experiments were compared to

examine the effect of different mixing states on the MTO product distributions. Figures 33 and

34 contain plots of the PFR and Berty reactor data collected for the 1.39% and 2.60% pre-coked

catalysts. Both of the figures are for the high conversion scenarios.

At low coke levels, as shown in Figure 33, both reactors produced more propylene than

ethylene, however the Berty reactor produced more ethylene than the PFR. Higher propylene

levels in the PFR are consistent with the results obtained in the initial and base case PFR runs.

Comparing the higher coked catalyst (Figure 34), there is clearly a difference in the distribution

of ethylene and propylene produced from different mixing states. On the more highly coked

catalyst, the production of ethylene is enhanced at higher partial pressures of oxygenates, which

can also be observed with the lower coked catalysts (Figure 32). This is similar to the increase

of olefin selectivity with increasing WHSV as reported in the literature [21], [22], [23].

However, there was no attempt to control the coke levels in these earlier studies, and it is likely

that the earlier data correspond to a more highly coked catalyst. The fact that the change of

ethylene:propylene ratios is not observed in the fresh catalyst (Figures 13-14, and 20) or at low

coke levels indicates that coke alters (or blocks) the reaction paths. In addition, the presence of

coke also causes a significant change in the following parallel reactions:

COH2OHCH 23 +→ (4)

84

Figure 33: Mole fraction of light olefins and saturates for the PFR (81.4% conversion, 1.46 wt% coke) and Berty reactor (81.8% conversion, 1.39 wt% coke). [Water, argon, dimethyl ether, methanol free basis]

0.00

0.10

0.20

0.30

0.40

0.50

0.60

BertyPFR

85

Figure 34: Mole fraction of light olefins and saturates for the PFR (62.1% conversion, 2.63 wt% coke) and Berty reactor (69.6% conversion, 2.60 wt% coke). [Water, argon, dimethyl ether, methanol free basis]

0.000.050.100.150.200.250.300.350.400.450.500.55

BertyPFR

86

222 HCOOHCO +⇔+ (5)

Neither of these reactions occurred in the Berty reactor, but Reaction (4) did take place in the

PFR and was therefore highly dependent on the methanol partial pressure. The fact that carbon

dioxide was not observed with a fresh catalyst (Figure 13-17 and 20), but increased with

increased coke content, indicates that Reaction (5) is enhanced by the presence of coke. In

addition, since the coke levels on the fresh catalyst in Figure 32 are between 1.39 wt% and 2.60

wt% (the coke levels from the Berty reactor), the type (or location) of coke also influences

Reaction (5). The presence of carbon dioxide in the Berty reactor product stream also indicates a

change has occurred to the catalyst. The only explanation for these observations is that the

coking process in the Berty reactor was such that certain active sites on the catalyst were

modified or blocked to produce the changes in product distribution just discussed. This in terms

would conclude that different mixing states does not directly change the light olefin distribution,

but does so indirectly by altering the reaction and mechanistic pathways of the catalyst through

the deposition of coke in a CSTR mixing environment.

87

CHAPTER 5

CONCLUSIONS AND RECOMMENDATIONS

The goal of the investigation presented within this thesis was to determine if different

types of mixing states would alter the MTO product distribution. Two reactors, a PFR and a

Berty reactor, were successfully setup to provide two different mixing states. It is uncertain if

the different mixing states directly influence the olefin product distribution.

However, significant changes in olefin concentrations were observed for ethylene and propylene

after running PFR experiments with spent catalyst from the Berty reactor. At total oxygenates

conversion less than 80%, the ethylene to propylene ratio that was less than 1.0 increased to a

value greater than 1.0 (there was actually more ethylene than propylene in the products). In

addition, this switch in the ethylene to propylene ratio was not observed in the PFR runs with

fresh catalyst. This indicates that the change in the olefin distribution was not a direct effect of

the mixing states, but rather due to the aging of the catalyst in the Berty reactor through coking,

which may have altered the reaction pathways of the MTO reaction. On the other hand, the

mixing state had a direct influence on the selectivity of saturates. Significantly higher saturate

concentrations were produced in the Berty reactor. Assuming that the saturates result from

cracking and/or oligomerization, this is consistent with a series nature for the reaction sequence

and the lower oxygenate concentrations observed in the Berty reactor.

Initial experiments with the PFR using fresh catalysts demonstrated that as conversion

decreased with time on stream, the DME concentration remained constant as the methanol

concentration increased. This indicates that there is more than one parallel reaction for methanol

to follow. Methanol could either (1), react directly to form olefins or (2), be converted to DME,

88

which would then form olefins via a hydrocarbon pool mechanism. The PFR runs using fresh

catalysts also indicated the formation of carbon monoxide, which could be a third reaction path

for methanol. However, the presence of carbon monoxide was only observed in the PFR,

indicating that carbon monoxide formation requires the high oxygenate concentrations provided

by the plug flow environment.

The TGA/MS analyses during coke burn-off indicated that there were different species of

coke deposited on the catalyst in the Berty reactor and the PFR. This was verified by the

presence of multiple carbon dioxide peaks as the temperature was ramped to 800ºC. In addition,

carbon dioxide produced via water-gas shift was measured in the product distribution only when

using spent catalyst from the Berty reactor in the PFR. These observations, in addition to those

discussed earlier with the ethylene:propylene ratios, indicate that not only is there more than one

species of coke deposited on the catalyst, but the type of coke and/or location of the coke

deposited depends on the type of mixing state used.

For future directions of this investigation, SAPO-34 should be used as the main catalyst.

SAPO-34 and its derivatives have proven to achieve excellent olefin selectivity and should be

used in a mixing state investigation to further clarify the mixing state effects. Using ZSM-5 was

sufficient to prove that mixing states can affect the product distribution, however, ZSM-5 is not

as selective towards light olefins as SAPO-34, and results in more by-products. The

investigation should also cover a larger range of catalyst coke levels and oxygenate conversions.

During the course of the investigation presented in this thesis, only two coke levels and two

conversions were examined.

In terms of recommendations for equipment, the accuracy achievable with a GC/MS type

instrument would be improved with a capillary column capable of better alcohol separations.

89

The method for calibrating methanol used in this investigation was sufficient, but the accuracy

and repeatability of the data would improve if the trailing peak effects could be eliminated. In

addition, it would speed up the process of calibrating the instrument.

Second, a more convenient backmix reactor is needed. The Berty reactor was fully

functional and was able to provide the necessary mixing states for data acquisition, but

maintenance and operation of the reactor was time consuming. On a practical time scale, only

two to three experiments could be run per week, slowing progress. Plus, the Berty reactor is very

large and some people may not have the strength to take the reactor apart for cleaning and

catalyst loading. The Berty reactor was also limited to a maximum operating temperature of

450ºC, which is a limitation if higher temperatures are to be investigated. A smaller, more

compact reactor would be the best solution, such as a CSTR with spinning catalyst baskets for

mixing of the gases. It should be constructed of material that is capable of operating at

temperatures up to 600ºC.

Finally, a more advanced methanol pump would improve the liquid feed to the system.

The syringe pump used in the research required different syringe sizes depending on the

flowrates to be used. If too large of a syringe was used with very small flowrates, vapor pulsing

could occur. A fluid pump capable of accommodating a wide range of flowrates and which

could maintain accuracy at very low flowrates would be ideal.

90

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[37] Derouane, E. G., Gilson, J. P., and Nagy, J. B., “Adsorption and Conversion of Ethylene on

H-ZSM-5 Zeolite Studied by 13C NMR Spectroscopy,” J. Molec. Catal., 10, 331-340 (1981).

[38] Chen, G., Liang, J., Wang, Q., Cai, G., Zhao, S., and Ying, M., in New Developments in

Zeolite Science and Technology, ed. Murakami, Y., Iijima, A., and Ward, J.W., Elsevier, Tokyo,

907 (1986).

[39] Walsh, D. E., and Rollmann, L. D., “Radiotracer Experiments on Carbon Formation in

Zeolites. II,” J. Catal., 56, 195-197 (1979).

[40] Magnoux, P., Cartraud, P., Mignard, S., and Guisnet, M., “Coking, Aging, and Regeneration

of Zeolites: III. Comparison of the Deactivation Modes of H-Mordenite, HZSM-5, and HY

during n-Heptane Cracking,” J. Catal., 106, 242-250 (1987).

[41] Langner, B. E., “Reactions of Methanol on Zeolites with Different Pore Structures,” Appl.

Catal., 2, 289-302 (1982).

[42] Ione, K. G., Echevskii, G. V., and Nosyrva, G. N., “Study of Stability and Selectivity of

Catalytic Action of ZSM-Type Zeolites in Methanol Transformation,” J. Catal., 85, 287-294

(1984).

96

[43] Bibby, D. M., Milestone, N. B., Patterson, J. E., and Aldridge, L. P., “Coke Formation in

Zeolite ZSM-5,” J. Catal., 97, 493-502 (1986).

[44] Ducarme, V., and Vedrine, J. C., “ZSM-5 and ZSM-11 Zeolites: Influence of Morphological

and Chemical Parameters on Catalytic Selectivity and Deactivation,” Appl. Catal., 17, 175-184

(1985).

[45] Kärger, J., Pfeifer, H., Jürgen, C., Bülow, M., Schlodder, H., Mostowicz, R., and Völter, J.,

“Controlled Coke Deposition on Zeolite ZSM-5 and Its Influence on Molecular Transport,”Appl.

Catal., 29, 21-30 (1987).

[46] Behrsing, T., Jaeger, H., and Sanders, J. V., “Coke Deposits on H-ZSM-5 Zeolite,” Appl.

Catal., 54, 289-302 (1989).

[47] Froment, G. F., De Meyer, J., and Derouane, E. G., “Deactivation of Zeolite Catalysts by

Coke Formation,” J. Catal., 124, 391-400 (1990).

[48] Gusinet, M., and Magnoux, P., “Review: Coking and Deactivation of Zeolites, Influence of

the Pore Structure,” Appl. Catal., 54, 1-27 (1989).

[49] Anthony, R., “Methanol Conversion to Olefins,” presented at: 1988 AIChE Spring Meeting,

New Orleans, LA, March 6-10, 1-14 (1988).

APPENDIX

98

A. GCD calibration

1) All species except methanol

To calibrate for the different compounds present in the MTO product distribution,

mixtures of each gas were prepared and analyzed with the GCD. The mixtures used for

calibration were 100/0, 75/25, 50/50, 25/75, and 0/100, each number representing a mole

percentage. Each calibration used ethylene as one of the components because it was one of the

most abundant olefins in the product stream. Three samples of each mixture were taken for

every calibration. The averages of these samples were plotted as peak intensity versus mole

percent for each gas. By inserting a linear trend line through the data and forcing the y-intercept

to equal zero, the slope of the trend line produced the value of the k constant with the following

equation:

Measured peak area (y) = “k” * Mol % of component (x)

Typical R2 values for these trend lines were ~1.0. The following list contains the k

Gas Ion k valuesCH4 15 453764CO 28 3140885

C2H4 28 2685758C2H6 28 4288517C3H8 28 2221317C3H6 41 5030461C4H8 41 9872041

n-C4H10 43 11523616CO2 44 7998629DME 45 6918574

99

values obtained from the calibration procedure:

Using ethylene for every calibration also aided with checking the system to make sure the

equipment was producing consistent results. At the end of each calibration, the extracted k value

was compared with the k values obtained in previously completed calibrations. If the ethylene k

values varied significantly, the calibration was repeated.

2) Code used to apply calibration for methanol

The methanol peaks obtained from the GCD scans would tail, making it difficult to

calculate methanol concentration with only one calibration constant. To compensate for the

tailing and be able to acquire accurate measurements of methanol in the product streams, the

following k constants obtained for different areas of methanol were programmed into the

Microsoft® Excel 2000 analysis worksheet using Boolean operators:

These k constants were applied with the following code used to program the worksheet (Note:

these are not the actual lines programmed into the cells. The code was written this way for

clarification in this thesis).

1 If MeOH area ≥ 0 and < 10000, then use k constant 581074, else, go to line 2 2 If MeOH area ≥ 10000 and < 30000, then use k constant 1779574, else, go to line 3 3 If MeOH area ≥ 30000 and < 100000, then use k constant 2663354 else, go to line 4 4 If MeOH area ≥ 100000, then use k constant 3426586, else, go to line 5 5 Error

k's3426586 for MeOH areas >= 1000002663354 for MeOH areas < 100000 and >= 300001779574 for MeOH areas < 30000 and >= 10000581074 for MeOH areas < 10000 and >= 0

100

B. Reactor(s) operations procedures

Equipment:

• Vaporizer unit designed for the MTO experiments

• Back pressure regulator assembly designed for MTO experiments (back pressure regulator

was purchased from Tescom Corporation, model # BB-36AL1VVA4)

• Hewlett-Packard model 1800A GCD

• Lindberg temperature controller

• Barnant Company model # 689-0000 temperature controller

• Lindberg model 55035A furnace

• Omega type-K thermocouples

• Omega model # 115KC thermocouple thermometer

• Tegam Model 871A hand held digital thermometer

• Cole Parmer model # 74900 syringe pump

• Sherwood Monojec syringes

• Hewlett-Packard model # 0101-0113 soap film flowmeter

• Matheson model # 8272-0422 mass flow controllers (calibrated for 200 sccm O2)

• Matheson model 8274 multiple flow controller

Materials:

• Liquid methanol (Fisher Scientific, HPLC Grade, CAS 67-56-1)

• ZSM-5, prepared with colloidal silica binder (ZSM-5 from Zeolyst, colloidal silica from

Aldrich)

• Cylinder of pre-purified helium

101

• Cylinder of argon

1) Initial equipment preparations: PFR

1. Remove quartz beads and any catalyst from reactor.

2. With a pair of pliers, pull out the wire mesh insert from the bottom half of the reactor.

3. Rinse the wire mesh with DI water and place in an oven to dry at ~120ºC.

4. Place quartz beads in a separate oven and ramp up to 600ºC at 15ºC/min. Hold for 1 hr, then

cool (this burns off any coke that may have deposited on the beads).

5. Rinse inside of reactor tube with DI water, then brush with a tube brush. Periodically rinse

the reactor tube and brush with DI water.

6. Place reactor tube in oven with wire mesh to dry (~120ºC).

7. When all reactor parts are dry, remove from oven and insert the wire mesh, closed end first,

into the lower half of the reactor until the end of the wire mesh with the opening is even with

the bottom opening of the PFR.

8. Fill the upper half of the PFR with quartz beads. Lightly tap side of PFR until ~0.5” of open

space is visible (space from the top of the quartz beads to the opening of the upper half of the

PFR.

9. Pack the gap with quartz wool, flush with the top of the upper half of the PFR.

10. Weight out the required amount of catalyst.

11. Pour the catalyst into the lower half of the PFR.

12. Pack ~0.5” of quartz wool against the catalyst to hold the catalyst bed in place.

13. Attach the PFR to the bulkheads mounted to the furnace frame.

14. Adjust the PFR thermocouple so that the end of the thermocouple is in the catalyst bed.

15. Manipulate the valves so that inert gas flows through the reactor from MFC #2.

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16. Close the valves on the backpressure regulator and leak test the reactor fittings.

17. Open the valves on the back pressure regulator and set the purge gas flowrate and pressure

(there should be ~2.0 psig on the GCD valve).

18. Set the purge gas flowrate for the vaporizer section.

19. Wrap any exposed process lines with heat tape and insulation.

20. Fill the condenser beaker with ice and water.

2) Making a run with the PFR

1. Turn all power supplies to the heat tapes to half power.

2. Set the furnace temperature to 250ºC.

3. After 0.5 hr, run a GCD scan and check for water or any other contamination.

4. Turn the power supplies for the heat tapes to the final settings.

5. Increase the temperature of the PFR furnace to 450ºC by 50ºC increments.

6. Once everything is up to temperature, take another GCD scan to check for any

contamination.

7. Set the actual flowrates to be used for the experiment. MFC #1 supplies the helium to the

vaporizer and MFC #2 supplies the argon for the vaporizer (MFC #3 can either be used as the

purge gas for the reactor, or as an additional supply of helium for the vaporizer, depending on

the helium flowrates required).

8. Set the syringe pump to the required methanol flowrate and start the methanol feed to the

vaporizer. Adjust the pressure of the vaporizer to ~7 psig.

9. Wait 10 to 15 minutes for the methanol vapor flowrate to stabilize.

10. Once all the flowrates have stabilized, manipulate the valves to start the methanol/He/Ar feed

to the reactor (note: if MFC#3 is being used as a purge gas only, shut it off immediately. If it

103

was initially being used as a purge gas, but is required to supply the argon trace to the

reactor, switch the valves so that the argon goes to the vaporizer inlet).

11. After 1 to 5 minutes, begin sampling.

12. Once the experiment is completed, shut off the syringe pump.

13. Manipulate the valves so that the vaporizer is cut off from the reactor and the only purge gas

is received from MFC #2 (or MFC #3, depending on the flow conditions).

14. Run a sample scan to make sure all the methanol and other products have been flushed out.

15. Turn off all power supplies and cool the reactor.

3) Initial equipment preparations: Berty Reactor

1. Remove 3-piece catalyst basket assembly from the lower reactor shell.

2. Disassemble the catalyst basket and clean each basket piece with a scrub pad and DI water.

3. Rinse the basket piece and dry with a paper towel (preferably the wipes that do not leave

fibers behind).

4. Apply liquid methanol to a paper towel and scrub each piece of the catalyst basket. This will

remove any remaining carbon deposits.

5. Rinse the catalyst basket pieces with DI water.

6. Place the catalyst basket piece in an oven at 120ºC and leave until dry.

7. Take a 1/8” gas line and connect it to an inert gas cylinder. Turn up the pressure so a

sufficient amount of gas is flow out the line (about 60 to 80 psig) and blow out the underside

of the impeller located in the lower shell of the reactor.

8. Take a paper towel soaked with DI water and wipe the inside chamber of the upper and lower

halves of the Berty reactor shell.

104

9. Dry the inside chamber walls with another towel, then wipe the chamber walls with a wipe

soaked in methanol to remove the rest of the coke deposits.

10. Wipe the chamber walls of both the reactor halves with a wipe soaked with DI water and

allow to air dry.

11. Take the stainless steel ring that is placed between the two copper gaskets and scrub it with

DI water and a scrubbing pad.

12. Rinse the steel ring with DI water, dry, then clean with methanol.

13. Give the steel ring a final rinse in DI water and place in the oven with the catalyst basket

pieces.

14. Take one each of the upper and lower copper gaskets used to seal the Berty reactor and sand

all surfaces with a fine grit sand paper to remove any machining marks (~ 400 grit sandpaper

or finer).

15. Rinse the gaskets in DI water and place in the oven to dry with the other pieces (take care not

to cut fingers on the sharp edges of the copper gaskets).

16. Locate the eight hex bolts used to seal the two reactor halves together. Take the bolts to the

fume hood and with a wire brush, brush of the dried anti-seize while holding the bolt over a

garbage can (this will create a great deal of dust and will collect on everything in sight).

17. Once the bolts are brushed, place on a table and begin coating each bolt with a film of anti-

seize (this step is very important to make sure that the bolts can be removed after the reactor

has been heated. Make sure to coat all of the thread with anti-seize).

18. Locate the lower thermocouple that is inserted into the catalyst basket. Brush the anti-seize

off of the treads on the thermocouple and recoat with anti-seize. Soak a wipe with methanol

and clean any anti-seize residue that may have deposited on the thermocouple sheath.

105

19. Remove all the Berty reactor pieces placed in the oven. These pieces will be very hot!

20. Once cool, place the largest piece of the catalyst in the lower half of the reactor shell. There

should be three slots in the shell that will accept the three large tabs on the catalyst basket.

The basket must face the right direction. There is a small hole on the side of the basket piece

(towards the bottom edge of the basket). This is where the thermocouple is inserted from

outside of the reactor and must be pointed towards the inlet and outlet lines.

21. Once the largest piece is in place, insert the smaller two pieces into the largest piece, lining

up the marks on the top edges of the basket pieces. This will help line up the holes for the

thermocouples.

22. Insert the lower thermocouple into its port and tighten with a 5/8 open end wrench. Make

sure the thermocouple has gone all the way through every hole to the inside of the inner most

catalyst basket piece.

23. Place the thicker copper gasket in the recessed lip on the lower reactor shell.

24. Make sure all power supplied to the mixer is off. Turn the magnetic mixer by hand to maker

sure there is nothing making contact with the impeller.

25. Place the thinner copper gasket over the raised lip on the upper reactor shell.

26. Place the stainless steel ring on top of the copper gasket on the upper reactor shell. The steel

ring will only fit on one side.

27. Use two small pieces of tape to hold the copper gasket and stainless steel ring in place.

28. Load the required amount of catalyst into the catalyst basket.

29. Place the upper shell of the Berty reactor onto the lower shell. There are two steel pins,

which extend from the top of the lower reactor shell. These pins help line up the upper

reactor shell.

106

30. Once the upper shell is in place, look in between the two reactor shells to make sure that the

gaskets are making firm contact with the stainless steel ring.

31. Install the eight hex bolts through the upper reactor shell and screw in as far as possible by

hand.

32. With a hex wrench, finish screwing in the bolts just until each bolt bottoms out.

33. Place a two-foot cheater bar (just a piece of steel pipe) onto the hex wrench and begin to

torque the hex bolts in place. What works is to torque the bolts half way by tightening them

in a pattern similar to tighten a wheel on an eight-bolt truck tire. Once all the bolts are half

way tight, repeat the process, tightening the bolts further.

34. Turn on MFC

35. MFC #1 and adjust to maximum flow on the dial (use helium).

36. Close all exit lines to the Berty reactor.

37. It the gasket is completely sealed, the MFC should shut all the way off (the particular reading

depends on the flow controller you use).

38. If a leak is detected, tighten the hex bolts again.

39. Once the Berty reactor is leak free, open the exit lines and set MFC #1 to an appropriate flow

for purging air out of the reactor.

40. Place the heating jacket onto the Berty reactor and plug into the temperature controller.

41. Connect all thermocouples on the Berty reactor and heating jacket to the appropriate pieces

of equipment.

42. Turn on the cooling water to the magnetic drive. The flowrate should be ~ 2.0 L/min.

43. Turn on the mixer motor and RPM readout gauge. Adjust mixer speed to 1800 RPM.

107

4) Making a run with the Berty reactor

1. Turn all power supplies to the heat tapes to half power.

2. Set the furnace temperature to 250ºC.

3. After 2.0 hr, run a GCD scan and check for water or any other contamination in the exit lines.

4. Turn the power supplies for the heat tapes to the final settings.

5. Increase the temperature of the Berty reactor’s heating jacket to 450ºC by 50ºC increments.

6. Once everything is up to temperature, take another GCD scan to check for any

contamination.

7. Set up the condenser.

8. Set the actual flowrates to be used for the experiment. MFC #1 supplies the helium to the

vaporizer and MFC #2 supplies the argon for the vaporizer (MFC #3 can either be used as the

purge gas for the reactor, or as an additional supply of helium for the vaporizer, depending on

the helium flowrates required).

9. Set the syringe pump to the required methanol flowrate and start the methanol feed to the

vaporizer. Adjust the pressure of the vaporizer to ~7 psig.

10. Wait 10 to 15 minutes for the methanol vapor flowrate to stabilize.

11. Once all the flowrates have stabilized, manipulate the valves to start the methanol/He/Ar feed

to the reactor (note: if MFC#3 is being used as a purge gas only, shut it off immediately. If it

was initially being used as a purge gas, but is required to supply the argon trace to the

reactor, switch the valves so that the argon goes to the vaporizer inlet).

12. After 1 to 5 minutes, begin sampling.

13. Once the experiment is completed, shut off the syringe pump.

108

14. Manipulate the valves so that the vaporizer is cut off from the reactor and the only purge gas

is received from MFC #2 (or MFC #3, depending on the flow conditions).

15. Run a sample scan to make sure all the MeOH and other products have been flushed out.

16. Turn off all power supplies and cool the reactor. (VERY IMPORTANT: cooling water must

stay on until the reactor cools to 150ºC. If the water is turned off before the reactor has

cooled, permanent damage could result to the magnetic mixer)

109

C. Procedure for preparing and operating the TGA/MS equipment

Equipment:

• Netzsch Incorporated TGA analysis system, model STA 409 PC

• Balzers Instruments Quadstar MS, model 422

• Small crucible for sample (supplied by Netzsch, item # 459.478)

• Tweezers

Materials:

• Spent sample of ZSM-5 catalyst (100.000-250.000 mg)

• Cylinder of ultra high purity (UHP) helium

• Cylinder or compressed air

1) Procedure: Taking a TGA/MS sample

1. Clean a sample crucible with DI water or an alcohol and heat in an oven at 200°C for 0.5 hr.

2. After 0.5 hr, remove the crucible and allow cooling to room temperature.

3. Open the TGA furnace by pressing and holding the “up” and “safety” buttons at the same

time.

4. Check to see if the “purge” gas valve located on the side of the furnace is closed. If not,

close it.

5. Place the clean sample crucible onto the sample holder inside the furnace.

6. Open the software labeled “STA 409 PC on COM1 TASC 414-5”.

7. From the top menu, select “Diagnosis”, then “Adjustment Test…”.

8. Close the furnace (“down” + “safety” buttons).

110

9. Turn on the “protective” gas (helium in this case). Let it flow until a steady mass

measurement is obtained.

10. Press the “Tare” button.

11. Open the furnace by pressing the “up” and “safety” buttons at the same time.

12. Load the ZSM-5 sample into the sample crucible (do this carefully to avoid excessive force

on the balance instrument).

13. Close the furnace (“down” + “safety” buttons).

14. Wait for the balance to stabilize and then record the mass of the catalyst sample.

15. Close the outlet gas valve on top of the furnace with the supplied tool and close the

“protective” gas valve.

16. Turn on the vacuum pump.

17. Open the vacuum pump valve mounted on the furnace.

18. Pump the furnace chamber down to ~0.9 bar, then wait for five minutes.

19. Close the vacuum pump valve (leave the vacuum pump on).

20. With the “protective” gas flow meter turned all the way off, open the “protective” gas valve

on the side of the furnace (in these experiments it will be ultra high purity helium for the

“protective” gas).

21. Set the line pressure of the helium tank to 10 psig, and then open the helium flow meter all

the way.

22. This should back fill the furnace chamber. Continue this step until the pressure reads ~0.05

bar.

23. When the furnace is at ~0.05 bar, close the “protective” gas valve and turn the helium flow

meter all the way off.

111

24. Open the vacuum valve.

25. Pump down the furnace to ~0.9 bar again, but this time hold for 15 minutes.

26. Repeat steps 17 through 20.

27. When the furnace pressure is ~0.05 bar, adjust the helium flow meter to the required helium

flow rate (see instrument manual for flow meter adjustments for different gases).

28. With the “purge” gas flow meter all the way off, open the “purge” gas valve on the side of

the furnace (in these experiments it will be compressed air).

29. Set the line pressure for the compressed air tank to 8 psig.

30. Adjust the “purge” gas flow meter to the required air flow (see instrument manual for flow

meter adjustments for different gases).

31. Minimize the “STA 409 PC on COM1 TASC 414-5” software.

32. Open the “Paraset” software for the Quadstar 422 MS.

33. On the top pull down menus, select “Measure”, “MID”, and then the folder with the ion

parameters (if a parameter file has not be made yet, open one of the default files).

34. Set up the parameter file so it will display the ions to be looked for (example: 17 and 18 for

water, 44 for CO2, 32 for O2, and 28 for N2).

35. Save any corrections to the parameter file and close the “Paraset” software.

36. Open the “Measure + Sequencer” software for the Quadstar 422 MS.

37. Select the “OK→Start” button (wait about a minute after pressing the button).

38. On the new screen that appears after Step 35, select the “Run trend scan” button.

39. A new screen should then appear where the parameter file and the data file names are ask for.

Using the “browse” button, select the parameter file created in Step 31. For the data file

name, input a new file name.

112

40. Select the “Input completed” button. The Quadstar 442 software should now indicate that it

is waiting for a start signal from the TGA equipment.

41. Return to the ““STA 409 PC on COM1 TASC 414-5” software and close the tare screen if it

is still open.

42. From the top menu, select “File”, “New”.

43. Fill in the required information selecting “Sample” as the file type, and then press the

“Continue” button.

44. Select Tcalzero.tcx for the temperature calibration file (this will give an error warning, just

select OK).

45. Select Senszero.exx for the sensor calibration file (this will give an error warning, just select

OK).

46. On the next screen, set up the temperature program to be used with the TGA experiment.

47. Select the “Continue” button.

48. Give a name for the TGA file.

49. A tare screen will now appear. If an initial standby temperature has been programmed into

the temperature program, press the “Start Standby” button and wait for the furnace

temperature to stabilize at the programmed temperature. If an initial standby temperature is

not programmed into the temperature profile, skip to Step 48.

50. Check to see that the “protective” and “purge” gases are stabilized. If so, press the “Tare”

button.

51. Press the “Start Measurement” button. The TGA will now trigger the MS to start recording

MS data and the TGA will continue with the rest of the programmed temperature profile.

52. After the TGA/MS analysis is completed, allow the furnace to at least 300°C before opening.

113

53. Open the furnace (“up” + “safety” buttons) and remove the sample crucible. Close the

furnace (“down” + “safety”).

54. After the furnace temperature has reached 100°C, turn off the “protective” and “purge”

gases.

2) Correcting a TGA/MS correction for buoyancy

Equipment:

Same as described in Part 1

Materials:

• Cylinder of ultra high purity (UHP) helium

• Cylinder or compressed air

Procedure:

1. Follow Steps 1 through 40 as described in Part 1, skipping Steps 9 through 12 (no sample is

loaded during a correction).

2. Fill in the required information selecting “Correction” as the file type, and then press the

“Continue” button.

3. Repeat Steps 42 through 52. Note: the temperature profile used for a correction MUST be

identical to the profile used during a regular TGA/MS analysis.

114

115

116

117

118

119

120

121

122

123

124

125

126

127

128

129

130

131

132

E. Carbon balance for PFR Run #11 [ZSM-5 (L)] (Normalized without water)

1) Begin by calculating the amount of carbon that is unaccounted for:

[Carbon Feed] – [Carbon in Exit Gas] = [Carbon unaccounted for]

minmol5.540x10

minmol1.797x10

minmol2.351x10 433 −−− =−

fordunaccounteCarbon23.7%100x

minmol2.351x10

minmol5.540x10

3

4

=−

−

2) Assuming that most of the large carbon species are C9’s (assume average in C8-C10 range), calculate the molar flowrate of unaccounted carbon species:

[Carbon unaccounted for] x [Ratio of C9 to 9 mols of carbon] = [Molar flowrate of C9]

s'Cminmol6.156x10

carbonofmols9Cmol

xcarbonminmol5.540x10 8

594 −− =

3) Using the data from the table for Sample 2, calculate the molar flowrate of products measured in the product stream:

[Exit gas flow] x [mol fraction of products in exit stream] = [Molar flowrate of products]

minmol1.053x10Ar)fractionmol.29420(1.000xflowgasexitCalculated

minmol1.492x10 33 −− =−

133

4) Calculate the total molar flowrate of products (measured plus unaccounted for due to normalization)

[Molar flowrate of measured products] + [Molar flowrate of C9] = [Total molar flowrate]

flowratemolarproducttotalminmol1.115x10

minmol6.156x10

minmol1.053x10 353 −−− =+

5) Determine the percentage of products not accounted for by normalizing the data:

[Molar flowrate of C9] / [Total molar flowrate] = [Estimated % of products unaccounted]

datathegnormalizinfromfordunaccounteproductstheof5.5%100x

minmol1.115x10

minmol6.156x10

3

5

=−

−

134