small-scale fuel processing: kinetic study of co

School of Industrial and Information Engineering

Master of Science in Chemical Engineering

Small-Scale Fuel Processing:

Kinetic Study of CO Preferential Oxidation

Supervisor: Prof. Alessandra BERETTA

Co-supervisors: Dott. Roberto BATISTA DA SILVA JR., PhD

Dott.ssa Veronica PIAZZA

Candidate:

Valeria COLOMBO 899051

Academic Year 2018-2019

ABSTRACT

The aim of this Thesis work is the study of the kinetics of CO preferential oxidation (CO

PrOx) on a noble metal-based commercial catalyst, within the MICROGEN30 project

funded by the Ministry of Economic Development. In the experimental phase, tests were

carried out on the catalyst at the laboratory scale, under different operating conditions.

In particular, the effect of each species on the kinetics was investigated by varying the

inlet concentrations, both in the presence and in the absence of hydrogen. Two different

reactor configurations were exploited: diluted packed bed reactors at different dilution

ratios, and an annular reactor operating at very high space velocity under quasi-

isothermal conditions. Data gathered during the experiments were then used to perform

a kinetic analysis. In particular, a previously developed 1-d heterogeneous model for the

annular reactor was properly modified through the introduction of convenient rate

expressions that incorporate the major kinetic dependencies. The obtained expressions

are suitable to simulate the CO PrOx reactor.

Keywords: CO PrOx, preferential oxidation, fuel processing, kinetics, catalytic processes.

ESTRATTO

Scopo della presente tesi è lo studio della cinetica dell’ossidazione preferenziale di CO

(CO PrOx) su un catalizzatore commerciale a base di metallo nobile, nell’ambito del

progetto MICROGEN30 finanziato dal Ministero dello Sviluppo Economico. Nella fase

sperimentale, sono stati condotti degli esperimenti su scala di laboratorio in differenti

condizioni operative. In particolare, è stato indagato l’effetto di ciascuna specie sulla

cinetica della reazione variandone la concentrazione in ingresso, sia in presenza, sia in

assenza di idrogeno. Sono state impiegate due diverse configurazioni reattoristiche:

reattori a letto impaccato caratterizzati da diversi rapporti di diluizione, e un reattore

anulare operante ad alta velocità spaziale in condizioni quasi isoterme. I dati raccolti nel

corso degli esperimenti sono stati quindi utilizzati per uno studio cinetico. In particolare,

un modello 1-d eterogeneo per il reattore anulare, sviluppato in precedenza, è stato

modificato mediante l’introduzione di opportune espressioni che incorporano le

maggiori dipendenze cinetiche. Tali espressioni risultano utilizzabili ai fini della

simulazione di un reattore di CO PrOx.

Parole chiave: CO PrOx, ossidazione preferenziale, fuel processing, cinetica, processi

catalitici.

vii

CONTENTS

List of figures ............................................................................................................................. xi

List of tables ............................................................................................................................. xiv

1 State of the art .................................................................................................................... 16

1.1 MICROGEN30 .......................................................................................................... 16

1.1.1 Description of the unit ......................................................................................... 16

1.2 Hydrogen production .............................................................................................. 17

1.2.1 Main processes for hydrogen production ......................................................... 17

1.2.2 Steam reforming ................................................................................................... 19

1.2.3 Steam reforming in micro-CHPs ........................................................................ 20

1.3 Water gas shift .......................................................................................................... 21

1.4 Preferential oxidation of CO ................................................................................... 23

1.4.1 Introduction .......................................................................................................... 23

1.4.2 Catalysts reported in the literature .................................................................... 25

1.4.3 Mechanism of CO preferential oxidation on PGMs ........................................ 27

2 Experimental methods ..................................................................................................... 30

2.1 Description of the rigs .............................................................................................. 30

2.1.1 Feed section ........................................................................................................... 30

2.1.2 Reaction section (FBR plant) ............................................................................... 32

2.1.3 Reaction section (annular reactor plant) ........................................................... 33

2.1.4 Analysis section .................................................................................................... 34

2.2 Experimental procedures ........................................................................................ 36

2.2.1 Start-up of the rig ................................................................................................. 36

2.2.2 Execution of the experiment ............................................................................... 37

2.2.3 Axial temperature profiles (annular reactor) ................................................... 39

viii

2.2.4 Shut-down of the rig ............................................................................................ 39

2.3 Catalyst characterization ......................................................................................... 40

2.3.1 Main features of the catalyst ............................................................................... 40

2.3.2 Catalytic granules preparation ........................................................................... 42

2.3.3 Slurry preparation ................................................................................................ 43

2.3.4 Dip coating ............................................................................................................ 45

2.4 Thermodynamics ...................................................................................................... 46

2.4.1 Introduction .......................................................................................................... 46

2.4.2 Minimization of Gibbs’ free energy ................................................................... 46

3 Experiments in diluted packed bed reactors ................................................................ 49

3.1 Introduction .............................................................................................................. 49

3.1.1 Choice of the packed bed reactor ....................................................................... 49

3.1.2 Reactors used in this work .................................................................................. 49

3.1.3 Operating conditions ........................................................................................... 51

3.1.4 Apparent deactivation of the catalyst ................................................................ 52

3.2 Diagnostic criteria for heat transport limitations ................................................. 53

3.2.1 Introduction .......................................................................................................... 53

3.2.2 Interphase transport ............................................................................................. 54

3.2.3 Interparticle transport .......................................................................................... 55

3.2.4 Estimation of the transport properties .............................................................. 57

3.3 Reactor history .......................................................................................................... 59

3.3.1 BED1 ........................................................................................................................ 59

3.3.2 BED2 ........................................................................................................................ 62

3.3.3 BED3 ........................................................................................................................ 67

4 Experiments in the annular reactor ................................................................................ 72

4.1 Introduction .............................................................................................................. 72

ix

4.1.1 The annular reactor .............................................................................................. 72

4.1.2 V1 ............................................................................................................................ 73

4.2 Experiments carried out in the presence of hydrogen ........................................ 74

4.2.1 Stabilization phenomena ..................................................................................... 74

4.2.2 Effect of the GHSV ............................................................................................... 76

4.2.3 Effect of yCO ......................................................................................................... 79

4.2.4 Effect of yO2 .......................................................................................................... 82

4.3 Experiments carried out in the absence of hydrogen .......................................... 84

4.3.1 Introduction .......................................................................................................... 84

4.3.2 Effect of the GHSV ............................................................................................... 84

4.3.3 Effect of yO2 ........................................................................................................... 86

4.3.4 Effect of yCO ......................................................................................................... 88

5 Kinetic study ...................................................................................................................... 90

5.1 Introduction .............................................................................................................. 90

5.1.1 Rate of the reaction ............................................................................................... 90

5.1.2 Kinetic analysis in differential regime ............................................................... 92

5.2 Mathematical model of the annular reactor ......................................................... 93

5.2.1 Introduction .......................................................................................................... 93

5.2.2 Equations of the model ........................................................................................ 93

5.2.3 Mass transfer resistances ..................................................................................... 95

5.2.4 Reaction rates ........................................................................................................ 95

5.3 Study of CO oxidation in the absence of hydrogen ............................................. 96

5.3.1 Introduction .......................................................................................................... 96

5.3.2 Differential analysis ............................................................................................. 97

5.3.3 Integration of CO oxidation into the model of the annular reactor .............. 99

5.4 Study of CO oxidation in the presence of hydrogen ......................................... 104

x

5.4.1 Preliminary considerations ............................................................................... 104

5.4.2 Choice of a reaction scheme .............................................................................. 109

5.4.3 Integration of CO oxidation into the model of the annular reactor ............ 110

5.4.4 Comparison with methanation ......................................................................... 115

Conclusions ............................................................................................................................. 117

Bibliography ............................................................................................................................ 120

xi

LIST OF FIGURES

Figure 1.1: A scheme of the micro-CHP system developed by ICI Caldaie (from [1]). .. 16

Figure 1.2: Fuel processing of solid, liquid and gaseous fuels for hydrogen production

(from [3]). ................................................................................................................................... 18

Figure 1.3: Examples of reformers. A: top-fired reformers. B: wall-fired reformers (from

[2]). .............................................................................................................................................. 20

Figure 1.4: Performances of different types of catalysts for PrOx in terms of CO

conversion and reaction temperature window (from [12]). ............................................... 25

Figure 2.1: Annular reactor plant. .......................................................................................... 30

Figure 2.2: Example of calibration curve. .............................................................................. 32

Figure 2.3: Example of tube. ................................................................................................... 33

Figure 2.4: Example of a chromatogram obtained for column A. From left to right: H2,

O2, N2, CO. ................................................................................................................................. 35

Figure 2.5: Brooks control unit. .............................................................................................. 37

Figure 2.6: Scheme of the thermocouples used for the measurement of the axial

temperature profiles................................................................................................................. 39

Figure 2.7: Catalytic pellets observed at the optical microscope. The pellet on the right

was cut in half for the measurement of the thickness. ........................................................ 40

Figure 2.8: Logarithmic differential pore volume distribution vs pore diameter, obtained

through MIP. In red: powder. In green: pellets. .................................................................. 42

Figure 2.9: Mortar and pestle. ................................................................................................. 42

Figure 2.10: Hydraulic press. .................................................................................................. 43

Figure 2.11: Ball milling. .......................................................................................................... 44

Figure 3.1: BED1. ....................................................................................................................... 50

Figure 3.2.: Conversion drift at 60 °C over a 150-minute time period (BED3).

GHSV=160000 NL/h/kg. Inlet composition: 40% H2, 1% CO, 1% O2. ................................ 52

Figure 3.3: Effect of the GHSV in BED1. Inlet composition: 40% H2, 1% CO, 1% O2. ...... 60

Figure 3.4: Conversion drift in BED2. GHSV=240000 NL/h/kg. Inlet composition: 40% H2,

1% CO, 1% O2. Injections from 6 up to 15 were carried out around 30 mins after the first

five ones, at 10-minute intervals. ........................................................................................... 62

Figure 3.5: Effect of the GHSV in BED2. Inlet composition: 40% H2, 1% CO, 1% O2. ...... 63

xii

Figure 3.6: Effect of oxygen concentration in BED2. GHSV=240000 NL/h/kg. ................. 65

Figure 3.7: Check on the presence of interphase and radial temperature gradients in BED2

(GHSV=240000 NL/h/kg, 40% H2, 1% CO, 1% O2). In blue: the term at the left hand side

of each criterion (see 3.2.2 and 3.2.3). In red: the threshold. .............................................. 66

Figure 3.8: Trend of CO conversion for the three reference tests carried out in BED3. Inlet

composition: 40% H2, 1% CO, 1% O2. GHSV=160000 NL/h/kg. ......................................... 67

Figure 3.9: Conversion drift in BED3 (inlet composition: 40% H2, 1% CO, 1% O2.

GHSV=160000 NL/h/kg). ......................................................................................................... 68

Figure 3.10: Comparison among BED1, BED2 and BED3. Left: comparison between BED1

(unconditioned) and BED2 (conditioned in hydrogen) at GHSV=80000 NL/h/kg. Right:

comparison between BED2 and BED3 at GHSV=160000 NL/h/kg. Inlet composition: 40%

H2, 1% CO, 1% O2. .................................................................................................................... 70

Figure 3.11: Check on the presence of interphase and radial temperature gradients in

BED3 (GHSV=160000 NL/h/kg, 40% H2, 1% CO, 1% O2). In blue: the term at the left hand

side of each criterion (see 3.2.2 and 3.2.3). In red: the threshold. ...................................... 70

Figure 4.1: Reference tests performed on V1. Inlet composition: 40% H2, 1% CO, 1% O2.

GHSV=500000 NL/h/kg. .......................................................................................................... 73

Figure 4.2: CO conversion drift at 100 °C, 90 °C and 80 °C on reactor V1. Inlet

composition: 40% H2, 1% CO, 1% O2. GHSV=500000 NL/h/kg. ......................................... 74

Figure 4.3: Trend of the outlet flow rate of CO for the three injections taken at each

temperature. GHSV=500000 NL/h/kg. Inlet concentration: 40% H2, 1% CO, 1% O2. ...... 75

Figure 4.4: Effect of the GHSV in V1. Inlet composition: 40% H2, 1% CO, 1% O2. .......... 77

Figure 4.5: Axial temperature profiles for the tests performed at 300000 and 1500000

NL/h/kg. ..................................................................................................................................... 78

Figure 4.6: Axial temperature difference between the catalytic bed and the oven for the

tests performed at 300000 and 1500000 NL/h/kg. ................................................................ 78

Figure 4.7: Effect of CO concentration on V1. GHSV=500000 NL/h/kg. ........................... 81

Figure 4.8: Effect of oxygen concentration on V1. GHSV=500000 NL/h/kg. .................... 83

Figure 4.9: Effect of the GHSV in the absence of hydrogen on V1 (squares). The results

are compared with the ones of the experiments carried out in the presence of hydrogen

(triangles). Inlet composition: 40% H2, 1% CO, 1% O2. ....................................................... 85

xiii

Figure 4.10: Effect of oxygen concentration on V1 in the absence of hydrogen.

GHSV=500000 NL/h/kg. .......................................................................................................... 87

Figure 4.11: Effect of CO concentration on V1 in the absence of hydrogen. GHSV=500000

NL/h/kg. ..................................................................................................................................... 89

Figure 5.1: Bilogarithmic plot for the data at varying CO concentration for the differential

analysis. ..................................................................................................................................... 97

Figure 5.2: Bilogarithmic plot for the data at varying O2 concentration for the differential

analysis. ..................................................................................................................................... 98

Figure 5.4: Results of the model for the tests in the absence of hydrogen: effect of the

GHSV. ...................................................................................................................................... 102

Figure 5.5: Results of the model for the tests in the absence of hydrogen: effect of CO

concentration. .......................................................................................................................... 102

Figure 5.6: Results of the model for the tests in the absence of hydrogen: effect of O2

concentration. .......................................................................................................................... 102

Figure 5.3: Comparison between the model and Prova 43 (GHSV=500000 NL/h/kg, 1%

CO, 1% O2, no hydrogen). .................................................................................................... 103

Figure 5.7: Trends of yCO2 and yH2O as a function of the percentage of CO, at three

different temperatures. .......................................................................................................... 104

Figure 5.8: Trends of yCO2 and yH2O as a function of the percentage of O2, at three

different temperatures. .......................................................................................................... 106

Figure 5.9: Arrhenius' plot for CO oxidation in the absence of hydrogen. .................... 107

Figure 5.10: Arrhenius' plot for CO oxidation in the presence of hydrogen. ................ 108

Figure 5.11: Arrhenius' plots for CO oxidation in the presence of hydrogen. Left: low

temperatures. Right: high temperatures. ............................................................................ 108

Figure 5.12: Results of the model. Effect of the GHSV. ..................................................... 112

Figure 5.13: Results of the model. Effect of CO concentration. ........................................ 113

Figure 5.14: Results of the model. Effect of O2 concentration. ......................................... 114

xiv

LIST OF TABLES

Table 2.1: Properties of the catalyst pellets. .......................................................................... 41

Table 2.2: Results of BET. ........................................................................................................ 41

Table 2.3: Results of mercury porosimetry. .......................................................................... 41

Table 2.4: Properties of tube V1. ............................................................................................. 45

Table 3.1: Packed bed reactors used in this Thesis work. ................................................... 49

Table 4.1: Tests carried out on V1. ......................................................................................... 73

Table 5.1: Data at varying CO concentration for the differential analysis in the absence

of hydrogen. .............................................................................................................................. 97

Table 5.2: Results of the differential analysis on the reaction order with respect to CO.

..................................................................................................................................................... 97

Table 5.3: Data at varying O2 concentration for the differential analysis in the absence of

hydrogen. ................................................................................................................................... 98

Table 5.4: Results of the differential analysis on the reaction order with respect to O2. 99

Table 5.5: Kinetic parameters for CO oxidation in the absence of hydrogen. ............... 101

Table 5.6: Kinetic parameters for CO oxidation in the presence of hydrogen. .............. 112

Table 5.7: Kinetic parameters for methanation (from [30]). ............................................. 115

Table 5.8: Comparison between the initial rates of methanation and PrOx. ................. 116

1 STATE OF THE ART

1.1 MICROGEN30

1.1.1 Description of the unit

MICROGEN30 is a project funded by the Ministry of Economic Development within the

Industria 2015 program, and led by ICI Caldaie. The project consists in developing and

operating a micro-combined heat and power (micro-CHP) system of small-medium scale

for residential applications, based on PEM (Proton Exchange Membrane) fuel cells. It can

generate 10-30 kW of electric power and 50 kW of thermal power [1].

Figure 1.1: A scheme of the micro-CHP system developed by ICI Caldaie (from [1]).

As schematically represented in Figure 1.1, natural gas and water are fed to the system.

Part of the gas is sent to a sulfur removal unit, then to the steam reforming stage; part of

it is sent to a burner, together with air and the PEM stack tail gas.

A first section consists of a train of catalytic stages for the conversion of the fuel into

syngas, and for its subsequent purification. Natural gas is converted into synthesis gas

inside a steam reforming unit, operating in the 650-800 °C temperature range. The

endothermic reaction is carried out in the presence of a catalyst, and heat is provided by

HYDROGEN PRODUCTION 17

means of the burner. Together with steam reforming, water gas shift always takes place

inside the system. Being it an exothermic reaction, at such high temperatures CO is

converted to CO2 only to a limited extent. Since carbon monoxide a poison for the anode

of the PEMFC, it must be removed up to a very small concentration (10 ppm). In order

to do so, a high temperature WGS stage, a low temperature WGS reactor and a final CO

PrOx unit are present.

Electric power is produced by means of four stacks of PEM fuel cells, fed with the

hydrogen exiting the first section and with air. The hot water supply is obtained by

means of water-water heat exchangers: heat is recovered from both the cooling systems

of the fuel cells, and in between the catalytic stages.

In the following sections, a brief description of each of the reactions taking place in the

catalytic stages will be provided, with a summary of the possible solutions which are

available, or under study, for small-scale applications. Particular attention will be given

to the reaction of preferential oxidation of CO.

1.2 HYDROGEN PRODUCTION

1.2.1 Main processes for hydrogen production

Hydrogen is one of the most important commodities in the chemical industry and in the

refinery sector. Since hydrogen is an energy carrier, but not an energy resource, it has to

be produced.

Synthesis gas (syngas) is the basis for most of the hydrogen produced [2]. It is a mixture

of hydrogen, carbon monoxide and carbon dioxide, and its applications also include the

syntheses of ammonia and methanol, oxo-synthesis and Fischer-Tropsch synthesis,

making it one of the most important intermediates in the chemical industry. The

production method depends on the raw materials – mainly natural gas, naphta, heavy

vacuum residue, and coal.

Unconventional fuels such as waste materials or biomass (which can be converted to

hydrogen both through gassification and biological processes [3]) have also gained

interest in the last years, but hydrocarbons still represent the main source for the

industrial-scale production of hydrogen. Hydrogen can be also produced from water,

18 STATE OF THE ART

mainly by electrolysis: however, this process is not as important as the fuel-based ones,

since the energy demand of electrolysis is in no way comparable to the one of

hydrocarbon-based processes [2].

Figure 1.2: Fuel processing of solid, liquid and gaseous fuels for hydrogen production (from [3]).

For the sake of brevity, only the main processes based on the reforming of hydrocarbons

will be shortly described here. Synthesis gas produced through one of these treatments

is then usually subjected to some aftertreatment, depending on the intended application

[2]. Three main techniques can be used to produce hydrogen from hydrocarbons: steam

reforming (SR), partial oxidation (POX, either catalytic or non-catalytic), and

autothermal reforming (ATR) [3].

Steam reforming requires no oxygen, it is operated at a lower temperature with respect

to POX and ATR and produces a reformate with a high H2/CO ratio: however, it is

characterized by the highest emissions and requires an external heat source. It is

described in more detail in the following paragraph.

Partial oxidation of hydrocarbons can be catalytic, or non-catalytic. In principle, it

consists in the reaction of hydrocarbons with an amount of oxygen which is insufficient

for total combustion. Non-catalytic POX is the most common process for heavy

feedstocks, even if virtually all hydrocarbon mixtures are suitable [2]. Catalysts can be

HYDROGEN PRODUCTION 19

used to lower the operating temperatures: in the case of natural gas, Ni and Rh-based

catalysts are typically used. Due to the exothermic nature of the oxidation reactions,

temperature is difficult to control [3].

Autothermal reforming is inbetween SR and POX: it adds steam to catalytic partial

oxidation. The heat required by the steam reforming reaction is provided through partial

oxidation, taking place in the thermal part of the reactor. Hence, there is no need for an

external heat source, simplifying the system and making it more flexible. However, the

oxygen to fuel ratio must be carefully controlled at all times, and in most cases an

expensive air separation unit is required.

1.2.2 Steam reforming

Steam reforming involves the reaction of steam with the fuel, in the presence of a

catalyst. The desired reaction is, in the case of methane:

CH4 + H2O ↔ CO + 3 H2

The reaction is strongly endothermic. In addition to steam reforming, the water-gas shift

reaction (slightly exothermic) also takes place in the system, producing some CO2:

CO + H2 ↔ CO2 + H2O

To obtain satisfying yields, the working temperature is around 800 °C. Despite the main

reaction being characterized by an increase in the number of moles, steam reforming is

usually carried out at pressures up to 30 atm, since the downstream processes are usually

performed under pressure. The steam to methane ratio is an important process

parameter, since not only does it influence the outlet composition, but it also prevents

coke formation [2]:

C + H2O ↔ CO + H2

Kinetically speaking, methane reforming can be described as a first-order reaction, no

matter the operating pressure. While at low temperatures the diffusion rate is much

higher than the reaction rate, at high temperatures pore diffusion has a strong impact on

the conversion [2]. Catalysts for steam reforming can be cathegorized into two main

groups: non-precious metals (typically Ni) and precious metals from Group VIII

elements (usually Pt or Rh) [3], very often promoted with alkali which are known for

increasing the activity of the catalyst and to facilitate coke gassification [4]. Due to severe

20 STATE OF THE ART

heat and mass transfer limitiations, the effectiveness factor for the catalyst in

conventional steam reformers is usually very low: thus, since rarely is the activity of the

catalyst a limiting factor, less expensive Ni-based catalysts are usually preferred [3]. The

catalyst is usually in the form of thick-walled, 16-mm diameter Raschig rings [2]. Other

common shapes include spoked wheels, gear wheels, or rings with several holes, and

are advantageous because of the low associated pressure drop. The catalyst can be

precipitated (higher activity, more prone to sintering) or impregnated (preferred due to

their higher mechanical resistance). Common supports include α-Al2O3 and MgO. Sulfur

is a strong poison for the catalyst and must be removed from the feed: this is usually

done by means of a zinc oxide desulfurization system.

The reaction is carried out industrially inside tubular reactors, built using special alloys.

Tubes are heated from the outside in fire-box-type units, both wall-fired and top- or

bottom-fired: burner geometry, flame length and the distance from the flame to the

reformer are all parameters that influence the homogeneity of heat transfer. Natural gas,

or low-sulfur containing hydrocarbons are employed as fuel.

Figure 1.3: Examples of reformers. A: top-fired reformers. B: wall-fired reformers (from [2]).

1.2.3 Steam reforming in micro-CHPs

The ideal fuel for a PEM fuel cell would be pure hydrogen. However, cost and technical

constraints make it difficult to store hydrogen in the necessary amount. Hence, hydrogen

gas is usually generated on site and on demand. Typical feedstocks are natural gas,

gasoline, LPG, and methanol [4].

WATER GAS SHIFT 21

Low-temperature PEM fuel cells usually exploit hydrogen produced by external

reforming with steam, air, or both. Depending on the operating conditions, an outlet

stream containing 3-10% CO is obtained. The main features of a catalyst for steam

reforming in fuel cell applications include high activity towards the fuel of choice,

resistance to poisoning, reduced start-up times, mechanical resistance and stability at

high temperatures under both steady-state and transient conditions. The support is

usually in the form of a ceramic or metallic monolith, foam, or some other structured

inert.

Frequently used catalysts for methane steam reforming include Rh, Pt and Ru [5]: the

advantage of precious metal catalysts is their high activity, durability, and low tendency

to both coking and sulfur poisoning.

The steam reforming unit of the MICROGEN30 fuel processor is characterized by three

main features [1]: the use of a precious metal catalyst; an annular reactor, packed with

catalyst particles diluted with highly conductive SiC beads, being heat transfer crucial

for this process [5]; a proprietary design of the burner, developed by ICI Caldaie.

1.3 WATER GAS SHIFT

For the majority of industrial processes, the carbon monoxide content in syngas as it is

produced from steam reforming is higher than required [2]. The water gas shift reaction

is typically exploited to remove this undesired amount of carbon monoxide, and the

same reaction is carried out in fuel processors.

Water gas shift is a catalytic reaction which converts CO and water into CO2 and

hydrogen:

CO + H2O ↔ CO2 + H2

The reaction is equimolar, thus the equilibrium conversion does not depend on the

operating pressure. On the contrary, the equilibrium composition does depend on the

temperature: since the reaction is slightly exothermic, the operating temperature should

be as low as possible, compatibly with the activity of the catalyst. The usual temperature

ranges are 300-510 °C for the high-temperature shift (HTS), and 180-270 °C for the low-

temperature shift (LTS). The upper temperature limit for HTS is dictated by the

22 STATE OF THE ART

resistance of the catalyst to sintering, while the lower limit for LTS is dictated both by

the poor activity of the catalyst and by the need of preventing water condensation and

the subsequent damage to the catalyst. The reaction is carried out industrially inside

multi-stage adiabatic reactors with inter-stage cooling. The two catalytic systems are set

in a series configuration: first the HTS and then, after an intermediate cooling, the LTS,

which is necessary to reduce the amount of CO from 2% to <0.5% [5].

Catalysts for HTS are usually Fe-Cr2O3-based, with Cu very often used as a promoter [2].

They are supplied in the oxidized condition, and reduced in situ. More recently, both Al

and Ce have been proposed as substitutes for Cr, which is active and stable, but also

highly toxic and thus leads to high disposal costs [6]. HTS catalysts have a lower activity

with respect to LTS catalysts, but they are quite resistant to impurities. Rapid

temperature and pressure changes must be avoided, since they lead to the disintegration

of the structure [2]. Catalysts for LTS include Cu-ZnO-Al2O3, where Cu represents the

active species. These catalysts were developed in more recent times, and have the

advantage of being active at lower temperatures. They are very sensitive to sulfur

poisoning: hence, a ZnO guard bed is present upstream of the LTS reactor. Moreover,

the reduced Cu species is pyrophoric [5] and the discharge process must be carried out

very carefully. Other categories of catalysts include ceria and noble-metal based

catalysts, carbon based catalysts and nanostructured catalysts [6].

Since the activity of the LTS is not too high at such low temperatures, typical space

velocities are about 1500-2000 h-1 [5]. Both LTS and HTS catalysts are characterized by

large volumes and hence large heat capacities, leading to a very slow response in

transient operations – an important limit for fuel cell-integrated fuel processors. Water

condensation, which is possible in the case of sudden stops, is also detrimental for the

catalyst. For these reasons, conventional base metal catalysts are not indicated for water

gas shift in fuel processing systems.

Water gas shift catalysts for fuel processing applications include Pt. Pt-containing

catalysts can be used at high temperatures, are highly effective (especially if used on a

monolith support) and show a zero-order kinetics with respect to CO, leading to a good

performance independently of the inlet concentration of CO. However, these catalysts

should be carefully designed to avoid the strongly exothermic methanation reaction.

PREFERENTIAL OXIDATION OF CO 23

Gold-containing catalysts have also been developed. No other precious metal has shown

promising activity towards WGS [5].

1.4 PREFERENTIAL OXIDATION OF CO

1.4.1 Introduction

Water gas shift allows achieving an outlet composition with a typical CO content of 0.5-

2% v/v [7]. CO is known to strongly adsorb on the Pt anode of the PEM fuel cell,

hindering the electro-catalytic oxidation of hydrogen [5] and causing irreversible

damage [7]. A way to partially recover the cell potential drop associated to the presence

of CO is the so-called air bleed [5], i.e. the addition of air to the reformate and the

subsequent oxidation of some of the chemisorbed CO. However, this technique has a

negative impact on the fuel cell operation, since it leads to the consumption of some of

the fuel, and to the dilution of the reformate. Therefore, the residual CO should be

removed upstream of the fuel cell as thoroughly as possible: concentrations of <10 ppmv

are to be reached.

Different processes have been developed for the removal of this residual amount of CO,

among them the pressure swing adsorption (PSA), and the employement of selective

hydrogen membranes: both methods require sufficiently high pressures to be effective

[5]. Among the catalytic methods for CO abatement, methanation and preferential

oxidation are the most widespread. Due to the large availability of highly selective

catalysts, efficient process control, lower operation costs and relatively simple

implementation [7], preferential oxidation is vastly employed in fuel cell applications.

Industrial applications of CO PrOx in hydrogen-rich streams date back as far as the early

1960s [5], when Engelhard developed a highly active and selective Pt-based catalyst,

named Selectoxo™, and a process for the selective oxidation of carbon monoxide in

ammonia synthesis gas. In particular, this process could be used for the treatment of gas

streams containing from a few parts per million up to 3% v/v of carbon monoxide, using

air in a range of about 3:1 to 0.25:1 oxygen to CO ratio, and an optimum GHSV around

5000 cubic feet per hour per cubic foot of catalyst [8].

24 STATE OF THE ART

Catalysts for CO PrOx are well-developed at the industrial scale: however, small-scale

fuel processors are associated to a number of constraints which lead to the development

of new catalysts [9].

A catalyst for preferential oxidation in fuel cell systems should satisfy the following

requirements [5]:

• lowering CO concentration down to <10 ppmv;

• showing high selectivity towards CO oxidation with respect to hydrogen

oxidation. Low selectivity is not only associated to an unnecessary consumption

of hydrogen to form water, but also to over-dilution of the reformate with

nitrogen as a result of excessive air injection;

• avoiding undesired side reactions. Due to the presence of large amounts of

hydrogen and CO2, the reverse water gas shift reaction (rWGS)

CO2 + H2 → CO + H2O

might occur, leading to CO production, especially at low space velocities and as

CO concentration approaches zero [10]. Being it an endothermic reaction, it is

favoured as the temperature increases. CO2 methanation

CO2 + 4 H2 → CH4 + 2 H2O

should also be prevented, since it leads to a large fuel consumption and to

runaway temperature excursions. It is especially favoured on Ru catalysts at

temperatures approaching 200 °C;

• operating within the range of temperatures and GHSVs of the fuel processor.

Thus, the inlet temperature should be compatible with the outlet one of the WGS

stage, or also conveniently to the one at which the fuel cell is operated (around

80-100 °C). At the same time, CO and hydrogen oxidation are highly exothermic

reactions: the heat released in the process can be recovered. The catalyst should

also show adequate activity within a wide range of space velocities, especially at

maximum flow.

• showing good chemical and mechanical resistance to the temperature cycling, air

exposure and water condensation associated to start-up and shut-down

procedures [9]. The catalyst must be stable even after thousands of start and stop

operations.


1.4.2 Catalysts reported in the literature

Catalysts for selective oxidation of carbon monoxide can be classified into three

categories [11]: noble metal-based, Au-based and transition-metal based catalysts.

Figure 1.4: Performances of different types of catalysts for PrOx in terms of CO conversion and reaction temperature

window (from [12]).

Noble metals active for CO PrOx include Pt, Pd, Ru, Rh and Ir, very often used in

combination with some kind of promoter or as a bimetallic nanostructure, and

supported on alumina or other oxides. Hulterberg et al. [10] compared the performances

of the previously mentioned five noble metals, either pure or supported on Ni and Co

oxides, in terms of carbon monoxide conversion and activity towards undesired

reactions, in particular the rWGS: among them, Ir was found to be the most active, with

an activity almost linearly increasing with the temperature. Rh displayed a similar

behaviour, but was proven to be less active than Ir: if mounted on Ni oxide, its

performances improved greatly. A Rh/MgO catalyst was also proposed by Han et al.

[13], owing to its low activity towards side reactions and high selectivity and activity at

temperatures up to 300 °C. Pt catalysts are probably the most extensively studied among

the noble metal-based ones: the activity of this species can be significantly increased by

using a Co oxide support [10]. However, Pt was also found to be very active towards

rWGS and tends to lose selectivity at higher temperatures. Ru, despite its relatively low

cost, presents a series of drawbacks [5]: it operates at temperatures well above the one of

the fuel cell, it is active towards methanation and deactivates both in presence of air and

liquid water. Therefore, special control strategies are required if this catalyst is used.

Mariño et al. [14] tested the performances of a ceria-zirconia-based Pd catalyst and found

26 STATE OF THE ART

it to have a very poor activity and selectivity towards CO oxidation if compared to the

corresponding Pt and Ir catalysts, possibly due to its oxidation to PdO.

The main drawback of PGMs is the poor activity in the low temperature range, since the

surface is predominantly covered with CO, hindering O2 adsorption [12]. Reducible

oxides active for oxygen storage, such as CeO2, have been proven to have a significant

impact on the performance of noble-metal based catalysts [14]. In general, reducible

oxides have the advantage of weakening CO adsorption while also providing additional

sites for oxygen adsorption or activation, turning the competitive Langmuir-

Hinshelwood mechanism between CO and oxygen for active sites into a non-competitive

one [15], [12]. Among reducible oxides, iron oxide has shown remarkable performances

[16], [17]. The performances of PGMs can be also improved by means of alkali metal

cations, such as Cs and Rb: if the low selectivity is assumed to be related to a spillover-

mediated hydrogen oxidation reaction, the enhancing effect of alkali can be explained

by the fact that these species are capable of supporting the hydroxyl groups required for

this unselective path [18].

Au-based catalysts have the advantage of showing good performance at lower

temperatures (around 100 °C) with respect to other noble metals, thus allowing a

straightforward implementation of the PrOx reactor in the same cooling circuit as the

PEM fuel cell, working at around 80 °C [19]. Despite being a poor catalyst in the pure

form due to the weak interaction with adsorbates, Au possesses high activity if highly

dispersed on a metal oxide support [14].

CuO-CeO2 are considered an economical alternative to noble-based metal catatalysts

[20]: these catalysts have been found to have good catalytic activity and selectivity.

Avgouropoulos et al. [21] compared the performances of Pt/γ-Al2O3, Au/α-Fe2O3 and

CuO–CeO2 catalysts and concluded that, while the Au-based catalyst is the most active

(capable of achieving 100% conversion at 45 °C for high enough contact times) and the

better one in the low temperature range, the CuO–CeO2 is the most selective and

preferable at higher temperatures; however, Pt/γ-Al2O3 is the most resistant towards CO2

deactivation, its performances in terms of CO oxidation being more or less unaffected

by CO2 partial pressure.


1.4.3 Mechanism of CO preferential oxidation on PGMs

Kahlich et al. [22] performed a series of experiments on Pt/Al2O3 in simulated reformed

gas and derived a simple power-law rate equation for CO oxidation through a

differential analysis of experimental data:

𝑟𝐶𝑂 = 𝑘𝑃𝐶𝑂0.42𝜆0.82

This expression, obtained through differential analysis, is consistent with the low rate

branch regime. In the so-called low rate branch, at low temperatures and/or λ =

2𝑃𝑂2 𝑃𝐶𝑂⁄ values, the surface is predominantly covered by CO. The low activity of PGM

catalysts at low temperatures can thus be explained by the presence of the CO adlayer,

which hinders the adsorption of oxygen (representing the RDS). As the temperature

increases, the desorption of CO becomes more and more favoured, and the adsorption

of oxygen is facilitated. Similar power-law expressions have also been proposed by other

authors [23].

Bissett et al. [24] derived rate expressions for CO and hydrogen oxidation over a Pt

catalyst by considering the following reaction mechanism:

CO + * ↔ CO*

H2 + 2 * ↔ 2 H*

O2 + 2 * → 2 O*

H* + O* → OH* + *

OH* + H* → H2O + *

CO* + O* → CO2 + *

and making a series of assumptions, mainly adsorption-desorption equilibrium for CO

and hydrogen, full CO coverage (θCO=1), negligible O2 desorption and CO adsorption

rate proportional to the square of the vacant sites. Under these hypotheses, the following

overall rates are obtained (𝑟𝐻2 is a linear combination of the two):

𝑟𝑂2 =𝑘𝑂2𝑘𝐶𝑂

𝑥𝑂2

√𝑥𝐶𝑂

𝑟𝐶𝑂 =2𝑘𝑂2𝑥𝑂2

√𝑥𝐻2 + 𝑘𝐶𝑂√𝑥𝐶𝑂

28 STATE OF THE ART

Choi and Stenger [11] proposed a reaction model for a Pt-Fe catalyst in which three

simultaneous reactions (CO oxidation, H2 oxidation and water gas shift) are considered:

−𝑟1 = 𝐴1exp(−33092

𝑅𝑇)𝑃𝐶𝑂

−0.1𝑃𝑂20.5

−𝑟2 = 𝐴2exp(−18742

𝑅𝑇)𝑃𝑂2

0.5

−𝑟3 = 𝐴3exp(−34104

𝑅𝑇)(𝑃𝐶𝑂𝑃𝐻2𝑂 −

𝑃𝐶𝑂2𝑃𝐻2𝐾𝑝(𝑇)

)

and emphasized the importance of accounting for the water gas shift reaction for a

reliable description of the reacting system.

Preferential oxidation on non-promoted platinum group metal (PGM) catalysts has been

long assumed to proceed through a simple, competitive Langmuir-Hinshelwood

mechanism between O2, CO and H2 [12], [15]. If a purely competitive reaction

mechanism is assumed for CO and H2, one should expect the presence of hydrogen to

have the only negative effect of lowering the selectivity towards CO oxidation [12]. This

is indeed the case for the high rate regime, characterized by a low CO surface coverage

and occurring at high temperatures, and/or high λ values: oxygen approaches total

conversion, and hydrogen and CO compete for it.

Still, a purely competitive mechanism has been excluded by some authors for Pt/Al2O3

catalysts [18]: such a mechanism should lead to a selectivity monotonically decreasing

up to zero as the amount of CO increases, but this contradicts the experimental

observation. The authors theorize a parallel, spillover-based oxidation pathway for

hydrogen.

Most importantly, hydrogen strongly enhances the reactivity of CO at low temperatures,

visibly lowering the CO light-off temperature [25]. On the opposite side, CO is known

to inhibit the ignition of H2, which would normally oxidize even at room temperature

[26].

Many hypotheses have been made to explain the promoting effect of hydrogen. It was

proposed that the heat of hydrogen oxidation increases the surface temperature,

promoting CO oxidation: however, this hypothesis is to be discarded, since it has been

proven that CO oxidation starts before the one of hydrogen [25]. Moreover, the extent of


the decrease in the light-off temperature related to the exotherm only is too limited, and

in any case not proportional to the concentration of hydrogen [27]. The simple

competition for active sites between hydrogen and CO and the related thermal effects

alone seem unable to describe the enhancement in the reactivity of CO.

Some authors [27] theorized a hydrogen-related reduction in the adsorption heat of CO:

this is in contradiction with other works proving no significant decrease in the surface

coverage of CO in the presence of hydrogen [28]. Other authors [22] theorized the

formation of formate species (consuming Pt-bonded CO) on the alumina support to

explain the increase in the reaction rate.

A mechanistic model which couples CO and H2 oxidation has been suggested by

Mhadeshwar and Vlachos [29]. It includes alternative, indirect routes for the oxidation

of CO such as the carboxyl-path, where hydroxyl reacts with adsorbed CO to form CO2

via a carboxyl intermediate [25]:

CO* + OH* ↔ COOH* + *

COOH* + * ↔ CO2* + H*

Formate-related routes are included in the model, as well, but have been suggested to

be negligible for PrOx [12]. Despite the low surface concentration of hydrogen-

containing species (H and OH) and the negligible conversion of H2 at the CO light-off

temperature, the carboxyl-path has been proposed by some authors [25] to be the main

reason for the promoting effect of H2 on Pt and Rh catalysts. The key step is the formation

of hydroxyl, which is able to react with CO at lower temperatures with respect to oxygen,

thanks to a lower activation barrier. Each adsorbed hydrogen atom oxidizing a CO

molecule is regenerated at the end of the cycle, and thus able to oxidize more CO

molecules before reacting with hydroxyl in a termination step.

In some cases, indirect oxidation of CO by OH has also been proposed to be the

dominant pathway, such as on a Ir-Fe catalyst [12]: the formation of the hydroxyl is in

this case the rate determining step for PrOx, and the oxidation by OH prevails on the

one by atomic O.

2 EXPERIMENTAL METHODS

2.1 DESCRIPTION OF THE RIGS

Two different rigs were used in this Thesis work: the first one for the tests with packed

bed reactors (see Chapter 3), the second one with the annular reactor (see Chapter 4).

Both can be divided into three parts: a feed, a reaction and an analysis section.

Figure 2.1: Annular reactor plant.

2.1.1 Feed section

Four gases have been used in the experiments: nitrogen, hydrogen, air and CO (as a

mixture of 20% CO and nitrogen). He and Ar were also fed to the plant, being used as

carrier gases by the gas chromatograph.

Hydrogen and air are provided from common cylinders in the lab basement, while the

CO and nitrogen mixture is stocked in a cylinder on the lab balcony. Nitrogen is stored

as a liquid inside a tank, outside of the building. Pressure reducers, mounted on the

walls, are required to bring the pressure of the gases from the one at which the gas is

stored (100-200 bar) up to 4-5 bar.

DESCRIPTION OF THE RIGS 31

The reactants are fed through four lines, each one equipped with shutoff valves, a mass

flow controller (MFC) and two manometers, one upstream and one downstream of the

mass flow controller. Filters are also present in order to remove any impurity entrained

in the gases.

In the case of the first rig, the following MFCs were employed:

• a 200 NmL/min MFC for the nitrogen line;

• a 700 NmL/min MFC for the hydrogen line;

• a 200 NmL/min MFC for the air line;

• a 50 NmL/min MFC for the CO line.

while in the case of the second rig, those were the MFCs of choice:

• a 700 NmL/min MFC for the nitrogen line;

• a 3 NL/min MFC for the hydrogen line;

• a 100 NmL/min MFC for the air line;

• a 50 NmL/min MFC for the CO line.

The mass flow controllers are connected to a four-channel Brooks control unit. In order

to select the proper opening, a calibration is to be carried out in advance. The opening

percentage is linked to the volumetric flow rate through an approximately linear

relationship:

𝑄 = 𝑚 ∙ %𝑂𝑃 + 𝑞

To calibrate a mass flow controller means to determine the values of 𝑚 and 𝑞. In practice,

the flow rate is measured for different values of the opening percentage through a bubble

flow meter, a graduated long tube with a rubber bulb at the bottom: the volumetric flow

rate is the volume crossed by the bubble per unit of time. In order to ensure a reliable

result, the same measurement is repeated at least five times. A linear regression on the

obtained data is used to determine the values of 𝑚 and 𝑞.


Figure 2.2: Example of calibration curve.

The temperature of the lines is regulated through heating tapes containing electrical

resistances, and set to 120 °C through a TIC. This is required to avoid water condensation

inside the lines. Aluminum sheets are wrapped around the lines to guarantee the

minimization of thermal dispersions. The temperature measurement is carried out

through thermocouples, put in direct contact with the lines.

2.1.2 Reaction section (FBR plant)

The reactants coming from the different lines converge into one single line. The gaseous

mixture is then split into two streams – one sent to the bypass line, the other one to the

reactor. The bypass line is used whenever the reactants are to be sent to the analysis

section without going through the reactor. Both lines are equipped with an upstream

shut-off valve: in the case of the reactor line, a downstream valve is also present to avoid

any reactant backflow. The streams coming from the two lines then mix again

downstream of the reactor.

With regard to the FBR plant, the reaction section includes the reactor, connected to the

heated lines through junctions and whose temperature is regulated through a vertical

Carbolite oven, with a height of 18 cm and an internal diameter of 12.5 mm. The bed

temperature is aquired through a K-type thermocouple, whose hot junction is located at

half of the bed height.

Three different fixed bed reactors were used during the experiments: however, they

were all very similar, differing only for the amount of catalyst and diluent. The reactor

y = 31,518x + 5,0593R² = 0,9999

0.000

0.050

0.100

0.150

0.200

0.250

0.300

0.350

0.400

0 2 4 6 8 10 12

Flo

w r

ate

[Nm

L/m

in]

Opening [%]


is a quartz tube with a 7-mm internal diameter, and a shrinkage in the bottom part: a

more accurate description is given in 3.1.2. In order to avoid any leakages, it is equipped

with a plastic cap and a punctured stopper, through which the thermocouple is inserted.

2.1.3 Reaction section (annular reactor plant)

In the case of the annular reactor plant, the reactor is located inside a horizontal,

cylindrical three-zone Carbolite TZF 12/38/400 furnace, with a length of 45 cm and an

internal diameter of 6 cm. The oven is heated by three independent resistors, each one

regulated through a PID controller which relies on a N-type thermocouple as measuring

element. The set point temperature is selected for the central portion of the oven: by

setting the temperature difference between this part and the lateral ones equal to zero,

the thermal uniformity of the oven is guaranteed.

One single annular reactor was used in the experiments. It consists of a 99.8%-pure

alumina tube, with a thin catalyst coating of known length and mass in its terminal

portion, coaxially inserted into a quartz tube. The reacting mixture flows through the

annular section included between the outer diameter of the alumina tube, and the inner

diameter of the quartz cylinder. A more detailed description is provided in 4.1.2.

Figure 2.3: Example of tube.


Two K-type thermocouples are also inserted into both the oven and the alumina tube.

By making them slide, it is possible to derive the axial temperature profiles along the

oven and the bed.

2.1.4 Analysis section

A gas chromatographer is required to detect the species exiting the reaction section and

their corresponding concentrations. The operating principle of such instrument is the

different affinity of the gaseous species with a stationary phase.

The gaseous mixture exiting the reactor is injected into the instrument together with an

inert gas, called carrier, and enters a long and thin glass capillary known as column. The

gas chromatographer is equipped with two different columns, each one characerized by

its own stationary phase and carrier gas, and capable of detecting different species. More

in particular:

• one column is equipped with molecular sieves and uses Ar as carrier. The

detected species are H2, O2, N2, CH4 and CO.

• the other column (Plot Q) uses He as carrier and is used to detect air+CO, CO2

and H2O.

The carrier gases are fed to the gas chromatographer at a given pressure and at any time.

The device is hardly ever switched off. When it is not being used, two different methods

can be selected:

• Spegnimento: the TCD is switched off, while the columns and the injector are set

to a temperature close to the ambient one;

• Condizionamento: the TCD is switched off, while the columns and the injector are

heated up to the maximum allowable temperature in order to remove any

undesired compound.

Each species leaves the columns at a different time, called retention time, which depends

on its affinity with the stationary phase, but also on the temperature at which the column

is operated. The time which is required for the analysis is chosen depending on the

largest retention time among the ones of the different species.

A TCD (thermal conductivity detector) is used to detect the components of the gaseous

mixture, separated by the chromatograph. The operating principle is the same as the one


of the Wheatstone bridge, a device containing four resistors subjected to a constant

thermal flux. Two branches of the bridge are brushed by the carrier gases, while the other

two are swept by the gaseous flow leaving the column. As a component other than the

carrier gas comes in contact with the resistor, its temperature changes due to the different

conductivity of the gaseous flow. This leads to a variation of the resistance and to a so-

called imbalance of the bridge. Such imbalance generates an electrical signal, which

allows the identification of any compound leaving the column. The non-ideality of the

gas mixture might influence the quality of the results.

One chromatogram is obtained for each column: such graph depicts the potential

difference generated by the TCD as a function of the time. Each peak corresponds to a

different detected species, which can be identified by its retention time. The area under

the peak is proportional to its concentration: however, the proportionality factor

depends on the species.

Figure 2.4: Example of a chromatogram obtained for column A. From left to right: H2, O2, N2, CO.

In order to convert the areas into concentrations, a calibration is to be carried out.

Calibrating the GC means to calculate the so-called response factors, referred to a species

whose inlet and outlet flow rates are known (usually nitrogen).

The response factor of the i-th species 𝛼𝑖 is defined as:

𝛼𝑖 =

𝑄𝑖𝑄𝑁2𝐴𝑖𝐴𝑁2


where 𝑄𝑖 is the flow rate of the i-th species and 𝐴𝑖 the area of the corresponding peak.

Thus, by definition, 𝛼𝑁2 is equal to 1.

In practice, a mixture of nitrogen and of the species under interest is sent to the gas

chromatographer. The flow rates of the two species are associated to a certain opening

percentage of the mass flow controller, and are assumed as known. Once the area of the

two peaks is known, the response factor of the species can be easily computed. It should

be noted that the response factor might vary depending on the fluxes.

Once the response factors have been determined, the volumetric flux of the i-th species

can be calculated starting from the one of nitrogen.

2.2 EXPERIMENTAL PROCEDURES

2.2.1 Start-up of the rig

The suction hood must be working as the experiment is started, in order to avoid any

gas leakage to the working environment. For safety, the room was also equipped with a

CO sensor. First of all, the cylinder containing the CO and nitrogen mixture on the lab

balcony is opened, and so are the valves upstream and downstream of the pressure

reducers on the walls. Shut-off valves are also opened at the entrance of the system.

Before proceeding, it is important to check for any leakages inside the reactor. In order

to do so, the valve upstream of the reactor is kept open, while the one downstream of

the reactor and the one on the bypass line are kept closed. Nitrogen is fed to the reactor

at a certain flow rate: once the pressure indicated by the manometers downstream of the

MFCs has reached a value close to 1 bar, the nitrogen feed is stopped. If no pressure drop

is observed for a sufficiently long amount of time (10-15 s), one can assume no leakage

is present and the experiment can start. After this check, the reactor is to be isolated

through the two shut-off valves. Only the bypass line is left open.

The electrical resistances heating the lines, the temperature controllers and the oven are

then switched on: the temperature of the oven is set to the desired set point. The TCDs

are also turned on, by choosing the proper analysis method for the gas chromatograph

(usually left on Spegnimento after each experiment).

EXPERIMENTAL PROCEDURES 37

Figure 2.5: Brooks control unit.

Once the temperature of the lines has reached the proper set point value, the reactants

can be sent through the bypass line to the analysis section. In order to do so, the proper

openings are to be selected on the Brooks control unit (Figure 2.5). The total flow rate,

which is required for the estimation of the molar flow rates of the single species starting

from the data aquired from the gas chromatograph, is measured by means of a

stopwatch through the bubble flow meter. The flow rates of the single reactants can also

be measured as an additional check.

Through the data acquired by the chromatograph, it is possible to check whether the

actual composition reflects the nominal one. If not, the openings of the MFCs have to be

adjusted. Once the concentrations of the reactants are close to the desired ones and the

oven has reached the nominal temperature, the reactor line can be opened and the

bypass line is closed.

2.2.2 Execution of the experiment

Once the oven has reached the desired temperature and is stable, the products can be

injected into the GC for the analysis. When collecting the data, three injections were

usually performed and their results averaged in order to minimize the experimental

error. More in particular, by defining:

𝑓𝑖 =𝐴𝑖𝐴𝑁2

where 𝐴𝑖 is the area of the peak corresponding to the i-th species, and indicating as 𝑓1, 𝑓2

and 𝑓3 the values of such ratio for each injection, the volumetric flow rate of species i can

be calculated as:


𝑄𝑖 = 𝛼𝑖𝑓1 + 𝑓2 + 𝑓3

3𝑄𝑁2

where 𝑄𝑁2 is the volumetric flow rate of nitrogen in NmL/min, which is assumed to be

constant throughout the whole experiment.

In the case of column B, there is no nitrogen peak. Instead, a macro-peak corresponding

to a CO, oxygen and nitrogen mixture is present. In order to quantify the flow rates of

the species detected by column B, the area of apparent nitrogen is calculated as:

𝐴𝑁2,𝑎𝑝𝑝𝐵 = 𝐴𝑚𝑎𝑐𝑟𝑜−𝑝𝑒𝑎𝑘

𝐵 ∙ 𝑥𝑁2

where 𝑥𝑁2 is the fraction of nitrogen in the macro-peak, calculated from the areas

obtained through column A (again, the average value of the three injections is taken):

𝑥𝑁2 =𝐴𝑁2𝐴

𝐴𝑁2𝐴 + 𝐴𝑂2

𝐴 + 𝐴𝐶𝑂𝐴

The molar flow in μmol/min and the molar fractions are given by:

�̇�𝑖 =𝑄𝑖

0.022414

𝑦𝑖 =�̇�𝑖

∑ �̇�𝑖𝑁𝐶𝑖=1

Starting from the composition, other quantities of interest can be calculated, such as the

conversion of CO and oxygen, and the selectivity of CO:

𝜒𝐶𝑂 =�̇�𝐶𝑂,𝑖𝑛 − �̇�𝐶𝑂

�̇�𝐶𝑂,𝑖𝑛

𝜒𝑂2 =�̇�𝑂2,𝑖𝑛 − �̇�𝑂2�̇�𝑂2,𝑖𝑛

𝑆𝐶𝑂2 =0.5(�̇�𝐶𝑂,𝑖𝑛 − �̇�𝐶𝑂)

�̇�𝑂2,𝑖𝑛 − �̇�𝑂2

In order to verify the quality of the experiments, carbon and oxygen balances were used

during the experiments. Such quantities are defined as the ratio between the carbon

(oxygen) atoms in the products divided by the converted carbon (oxygen) atoms, i.e.:

𝐵𝐶 =∑ �̇�𝑝𝑟𝑜𝑑,𝑖 ∙ 𝑛𝐶,𝑝𝑟𝑜𝑑,𝑖𝑁𝑃𝑖=1

∑ (�̇�𝑟𝑒𝑎𝑐𝑡𝑗,𝑖𝑛 − �̇�𝑟𝑒𝑎𝑐𝑡,𝑗) ∙ 𝑛𝐶,𝑟𝑒𝑎𝑐𝑡,𝑗𝑁𝑅𝑗=1

=�̇�𝐶𝑂2 ∙ 𝑛𝐶,𝐶𝑂2

(�̇�𝐶𝑂,𝑖𝑛 − �̇�𝐶𝑂) ∙ 𝑛𝐶,𝐶𝑂

EXPERIMENTAL PROCEDURES 39

𝐵𝑂 =∑ �̇�𝑝𝑟𝑜𝑑,𝑖 ∙ 𝑛𝑂,𝑝𝑟𝑜𝑑,𝑖𝑁𝑃𝑖=1

∑ (�̇�𝑟𝑒𝑎𝑐𝑡𝑗,𝑖𝑛 − �̇�𝑟𝑒𝑎𝑐𝑡,𝑗) ∙ 𝑛𝑂,𝑟𝑒𝑎𝑐𝑡,𝑗𝑁𝑅𝑗=1

=�̇�𝐶𝑂2 ∙ 𝑛𝐶𝑂2 + �̇�𝐻2𝑂 ∙ 𝑛𝐻2𝑂

(�̇�𝑂2,𝑖𝑛 − �̇�𝑂2) ∙ 𝑛𝑂,𝑂2

These quantities should be of course very close to 1. A balance greater than one means

that the amount of products is being overestimated, while a balance smaller than one

indicates its underestimation. The formation of parasitic species, such as solid carbon,

might also alter the value of the balances.

Once the data related to a certain nominal temperature have been collected, the set point

temperature of the oven can be increased up to the next desired value. The usual

temperature range used for the experiments was 100 °C-300 °C, at 20 °C intervals.

2.2.3 Axial temperature profiles (annular reactor)

In the case of the annular reactor, the axial temperature profile is also taken for both the

oven and the reactor – not only along the length the catalytic bed, but up to 1 cm

upstream and downstream of it. This is done by letting the thermocouples slide inside

the oven and the alumina tube two millimiters at a time, and by taking the corresponding

temperatures (indicated on the TIC). This measurement is of great importance, since the

reactor can be deemed isothermal only as long as the axial temperature gradient does

not exceed 5 °C/cm. The results obtained in the experiments are also indicated on the

plots not as a function of the nominal temperature, but as a function of the average

temperature of the bed.

Figure 2.6: Scheme of the thermocouples used for the measurement of the axial temperature profiles.

2.2.4 Shut-down of the rig

At the end of the experiment, the set point of the oven is set to zero. As the temperature

starts to decrease, the oven can be switched off. The mass flow controllers and shut-off

valves associated to all the reactants but nitrogen are then closed: since both air and a

large excess of hydrogen are present, it is important to close air first in order to avoid the


formation of a flammable mixture. Nitrogen is left flowing inside the reactor in order to

clean it: after a while, it is possible to open the bypass line and isolate the reactor by

closing the shut-off valves on the reactor line. Nitrogen can then be closed, and so can

the inlet valves, the valves associated to the pressure reducers and the CO-nitrogen

cylinder outside of the laboratory.

Finally, the heating tapes are turned off through the TIC and the GC is set to the

Spegnimento method.

2.3 CATALYST CHARACTERIZATION

2.3.1 Main features of the catalyst

The catalyst used in the experiments of this Thesis work is a commercial one, provided

by ICI Caldaie S.p.A. in the form of spherical pellets. It was never used as such: it was

used before in the form of a powder in fixed bed reactors, then as a deposited slurry in

the annular reactor. The catalyst is based on a platinum group metal (PGM), presumably

supported on alumina, and is said to require no pre-activation on the product sheet.

Figure 2.7: Catalytic pellets observed at the optical microscope. The pellet on the right was cut in half for the

measurement of the thickness.

The average diameter of the spheres and the thickness of the active phase were evaluated

through the observation of a selected number of spheres at the optical microscope. In a

first set of measurements, four spheres at a time were observed and photographed. The

cross section of the spheres was then estimated starting from the photographs, by

knowing the enlargement scale (60x). The average external diameter could be then

calculated from the cross section. In a second set of observations, the same procedure

was carried out, but the spheres were before cut in half by means of sharp blades. This

CATALYST CHARACTERIZATION 41

time, the thickness of the active phase could be assessed, again by knowing the

enlargement scale (200x). Further details are reported in [30].

The average mass of the spheres was also estimated by means of a precision scale.

Starting from the mass and the diameter of the pellets, the density of the catalyst can be

simply calculated as:

𝜌𝑐𝑎𝑡 =𝑚𝑐𝑎𝑡

𝜋6𝑑𝑠𝑝ℎ𝑒𝑟𝑒3

The properties of the catalyst are reported in the following table:

Average external

diameter [mm]

Thickness of the

active phase [mm] Mass [kg] Density [kg/m3]

2.014 0.095 4.31∙10−6 1007.3

Table 2.1: Properties of the catalyst pellets.

Both the pellets and the powders were also characterized from a morphological point of

view by means of a BET analysis (Micromeritics TriStar 3000) and of mercury intrusion

porosimetry (Micromeritics AutoPore V). The first analysis had the aim of quantifying

the surface area of the sample, while the second one was used to quantify the porosity,

the pore volume and the pore size distribution. The results of the analyses are reported

in the following tables.

From the BET:

Spheres Powder

BET surface area

[m2/g] 175.4 176.0

Table 2.2: Results of BET.

From Hg porosimetry:

Spheres Powder

Porosity [%] 69.5% 82.7%

Total pore area

[m2/g] 307.5 245.7

Average pore

diameter [Å] 83.3 197.5

Table 2.3: Results of mercury porosimetry.


Figure 2.8: Logarithmic differential pore volume distribution vs pore diameter, obtained through MIP.

In red: powder. In green: pellets.

2.3.2 Catalytic granules preparation

In all the experiments carried out in the fixed bed reactors, the catalyst was used in a

granular form. In order to ensure a sufficiently homogeneous distribution of the active

phase and a uniform size distribution, a proper procedure is to be followed when

preparing the powder.

Figure 2.9: Mortar and pestle.

The catalytic spheres are ground by means of mortar and pestle. The product is then

sieved, in order to separate the finest particles (the ones passing through MESH 200)

from the bigger ones, which are again put in the mortar and pestled. Due to the very

small dimension of these particles, a uniform active phase distribution can be assumed.

However, such small size of the granules might also lead to unacceptable pressure drops

inside the fixed bed. Hence, it is necessary to increase the particle dimensions.


Figure 2.10: Hydraulic press.

The fines are first compacted in a tablet-making machine. The particles are put between

two metal disks into a hollow cylinder, which is then set on a hydraulic press. The

pressure applied through the press pushes the disks and compacts the fines in a tablet-

like shape. Finally, the tablets are ground and sieved. This time, larger particles are

collected (MESH 140-200, between 0.074 and 0.105 mm).

The same procedure is followed for the preparation of the powders which are required

for the slurry.

2.3.3 Slurry preparation

In the case of the annular reactor, the catalyst is present on an alumina tube in the form

of a thin coating, obtained from the deposition of a slurry. The slurry is a dispersion of

alumina powders in water: these were obtained from the commerical catalyst pellets

through the same procedure described in 2.3.2. For the slurry to be stable, a strong acid

is also required: nitric acid was used in the preparation. It works by charging the surface

of the particles, but it is also consumed in an oxide dissolution reaction [31].

For the subsequent deposition step to be effective, the rheological properties of the slurry

(in particular its viscosity) should be adequate. These are strongly influenced by both

the H2O/powder and the HNO3/powder ratios. Initially, ratios of 1.4 mL of distilled

water per gram of powder and 1.7 mmol of HNO3 per gram of powder were chosen.


Indeed, this was found to be the optimal composition in previous works with Rh-

impregnated powders [32].

Starting from a given mass of powders (in this case, 5.4125 g), the corresponding amount

of HNO3 is given by (in grams):

𝑚𝐻𝑁𝑂3 =1.7

1000∙ 𝑚𝑝𝑜𝑤𝑑𝑒𝑟𝑠 ∙ 𝑀𝑊𝐻𝑁𝑂3

Nitric acid was used as a 65% w/w concentrated solution. Thus,

𝑚𝑠𝑜𝑙𝑢𝑡𝑖𝑜𝑛 =𝑚𝐻𝑁𝑂3

0.65

Finally, the required amount of water to be added is given by:

𝑚𝑤𝑎𝑡𝑒𝑟 =1.4

1000∙ 𝑚𝑝𝑜𝑤𝑑𝑒𝑟𝑠 ∙ 𝜌𝑤𝑎𝑡𝑒𝑟 − (1 − 0.65) ∙ 𝑚𝑠𝑜𝑙𝑢𝑡𝑖𝑜𝑛

During the dip-coating procedure, it was observed that the preparation dried on the

alumina tube without sticking to it. Hence, it was necessary to add a further amount of

water to the mixture (≃4 g) in order for it to acquire the proper viscosity.

The slurry is then put in a plastic container with zirconia balls, whose mass is 8 times the

one of the powders, and subjected to a 24-hour long ball milling procedure: the plastic

container is put inside a metallic cylinder, then set on two counter-rotating cylinders,

revolving at a constant speed of 40 rpm and driven by an electrical motor. This step is

required to obtain a homogeneous dispersion.

Figure 2.11: Ball milling.


2.3.4 Dip coating

For the deposition of the catalytic layer on the alumina tube, a procedure named dip

coating is followed. Aim of this step is to obtain a uniform, well-adherent layer of catalyst

on the alumina tube.

For this reason, before proceeding, the surface over which the catalyst is to be deposited

is covered in primer: this treatment results in an increase of the superficial roughness,

and favours the adherence of the deposited layers. In order to do so, the alumina tube is

immersed into a graduated cylinder containing the primer, then dried for 30 minutes at

ambient temperature. Teflon tape is used to isolate the catalytic bed from the rest of the

tube. After this preliminary operation, the actual dip coating procedure can be carried

out.

The slurry is put inside a graduated cylinder. Then, the alumina tube is attached to a

vertical slide, and carefully immersed into the cylinder at a constant speed. After being

left inside the solution for a while, the tube is pulled up, again at a costant speed. The

withdrawal velocity has an effect on the thickness of the coating. To remove the solvent,

the tube is dried in the oven for 10 minutes, at 280 °C. This procedure is called flash drying

and guarantees a fast evaporation and good adhesion of the catalyst layer.

After both the deposition of the primer and the one of the catalyst, the tube is weighted

by means of a precision scale, repeating the measurement at least three times. At the end

of the procedure, the catalyst load can be calculated as:

𝑚𝑐𝑎𝑡 = 𝑚𝑡𝑢𝑏𝑒+𝑐𝑎𝑡+𝑝𝑟𝑖𝑚𝑒𝑟 −𝑚𝑡𝑢𝑏𝑒+𝑝𝑟𝑖𝑚𝑒𝑟

One tube was prepared and used for this Thesis work. Here are reported its properties.

Name

Mass of catalyst

[mg]

Length of the

catalytic bed [cm]

V1 14.3 2

Table 2.4: Properties of tube V1.


2.4 THERMODYNAMICS

2.4.1 Introduction

The main reactions which can take place inside a PrOx reactor are CO and hydrogen

oxidation: oxidations are typically strongly exothermic reactions, for which the

equilibrium conversion is expected to decrease at increasing temperatures. Water gas

shift, which is also exothermic, can in principle take place in the system: in particular,

the reaction is thermodynamically favoured in the whole range of temperatures used in

the experiments, with an equilibrium constant decreasing as the temperature increases.

Methanation might also in theory take place in the system: however, the formation of

methane was never observed. For this reason, this species was not taken into account

while calculating the equilibrium state.

During the experiments, the conversion of CO went quite often above the equilibrium

value at higher temperatures. While this might seem a violation of the thermodynamic

constraint at first, it is actually related to the occurrence of the fast, non-equilibrated

oxidation reactions. The decrease in the conversion of CO at high temperatures can be

explained through the reverse water gas shift reaction, converting part of the produced

CO2 into CO. The reaction is slower than the oxidations, and takes place only as all the

oxygen has been consumed.

For each test, the equilibrium composition was calculated and inserted into the graphs

as a dashed line. In particular, STANJAN Chemical Equilibrium Calculator [33] was

used to evaluate the equilibrium state of the mixture, treated as an ideal mixture of ideal

gases. It is worth noticing that the code does not consider liquid and solid species in the

calculations. STANJAN finds the equilibrium composition by minimizing the Gibbs’ free

energy of the mixture, using temperature, pressure and the inlet composition as inputs

for the calculation. Carbon dioxide and water were considered as the only possible

products.

2.4.2 Minimization of Gibbs’ free energy

Determining the equilibrium state of a system subject to the constraint of constant

temperature and pressure means minimizing the Gibbs’ free energy of the mixture [34].

Hence, its differential must be zero:

THERMODYNAMICS 47

𝑑𝐺𝑇,𝑃 =∑𝜇𝑖𝑑𝑛𝑖

𝑁𝐶

𝑖=1

= 0 (2.1)

For the sake of brevity, a system with

𝑁𝑅 = 𝑁𝐶 − 𝑟𝑎𝑛𝑘(𝑨) (2.2)

where 𝑁𝐶 is the number of components and 𝑨 is the matrix with an atomic species for

each row and a molecular species for each column, will be considered. However, the

approach can be simply generalized to complex systems.

Due to the stoichiometric constraint, the ratio between the variation in the number of

moles of a certain species and its stoichiometric coefficient in a reaction must be the same

for all the species taking part in such reaction. Hence, by defining

𝑑𝜆 =𝑑𝑛𝑖𝜈𝑖

(2.3)

as the extent of the reaction, it follows that:

∑𝜇𝑖𝑑𝑛𝑖

𝑁𝐶

𝑖=1

=∑𝜇𝑖𝜈𝑖𝑑𝜆

𝑁𝐶

𝑖=1

= 0 (2.4)

Since 𝑑𝜆 is arbitrary, the equation can be rewritten as:

∑𝜇𝑖𝜈𝑖

𝑁𝐶

𝑖=1

= 0 (2.5)

𝜇𝑖 is the chemical potential of species i. By exploiting Lewis’ definition of fugacity:

𝑑𝜇𝑖,𝑇 = 𝑅𝑇𝑑𝑙𝑛𝑓𝑖 (2.6)

and integrating between a reference state and a generic one, it follows that:

𝜇𝑖(𝑇, 𝑃, 𝒚) = 𝜇𝑖(𝑇, 𝑃𝑟𝑒𝑓) + 𝑅𝑇𝑙𝑛 (

𝑓𝑖(𝑇, 𝑃, 𝒚)

𝑓𝑖(𝑇, 𝑃𝑟𝑒𝑓))

𝜇𝑖(𝑇, 𝑃, 𝒚) = 𝜇𝑖(𝑇, 𝑃𝑟𝑒𝑓) + 𝑅𝑇𝑙𝑛(𝑎𝑖(𝑇, 𝑃, 𝑃𝑟𝑒𝑓 , 𝒚))

(2.7)

where 𝑎𝑖 is the activity of the i-th species. The fugacity of a species i in a mixure can be

expressed as:

𝑓𝑖(𝑇, 𝑃, 𝒚) = 𝑃�̂�𝑖(𝑇, 𝑃, 𝒚)𝑦𝑖 (2.8)


For a gas, the reference state which is considered in the estimation of the activity, is an

ideal pure gas at 1 bar. Hence,

𝑎𝑖(𝑇, 𝑃, 𝑃𝑟𝑒𝑓 , 𝒚) =𝑃𝑦𝑖

1[𝑏𝑎𝑟]= 𝑃𝑦𝑖 (2.9)

Since the mixture is treated as an ideal mixure of ideal gases (only gaseous species are

present, and the working pressure is far from the critical pressures of the species), it

follows that:

𝑓𝑖(𝑇, 𝑃, 𝒚) = 𝑃�̂�𝑖(𝑇, 𝑃, 𝒚)𝑦𝑖 = 𝑃𝑦𝑖 (2.10)

Going back to equation 2.5,

𝑑𝐺𝑇,𝑃 =∑𝜈𝑖𝜇𝑖(𝑇, 𝑃𝑟𝑒𝑓)

𝑁𝐶

𝑖=1

+ 𝑅𝑇∑𝜈𝑖𝑙𝑛(𝑎𝑖(𝑇, 𝑃, 𝑃𝑟𝑒𝑓 , 𝒚))

𝑁𝐶

𝑖=1

= 0 (2.11)

Re-organising,

∑𝜈𝑖𝜇𝑖(𝑇, 𝑃𝑟𝑒𝑓)

𝑁𝐶

𝑖=1

+ 𝑅𝑇∏(𝑎𝑖(𝑇, 𝑃, 𝑃𝑟𝑒𝑓 , 𝒚))𝜈𝑖

𝑁𝐶

𝑖=1

= 0 (2.12)

It is worth remarking that the reference state for the two terms must be the same. Since

the chemical potential of a pure compound is equal to its molar Gibbs’ free energy, the

first term can be re-written as:

∑𝜈𝑖𝜇𝑖(𝑇, 𝑃𝑟𝑒𝑓)

𝑁𝐶

𝑖=1

=∑𝜈𝑖𝑔𝑖(𝑇, 𝑃𝑟𝑒𝑓)

𝑁𝐶

𝑖=1

= 𝛥𝐺°𝑅(𝑇, 𝑃𝑟𝑒𝑓) (2.13)

where 𝛥𝐺°𝑅 is the variation of Gibbs’ free energy associated to the complete

transformation of a stoichiometric amount of reactants into products, at reference

conditions.

By going back to equation 2.12 the second term is usually expressed as:

∏(𝑎𝑖(𝑇, 𝑃, 𝑃𝑟𝑒𝑓 , 𝒚))𝜈𝑖

𝑁𝐶

𝑖=1

= 𝐾𝑒𝑞 (2.14)

where 𝐾𝑒𝑞 is the equilibrium constant for the reaction. Thus, relation 2.12 becomes:

𝐾𝑒𝑞 = exp (−𝛥𝐺°𝑅(𝑇, 𝑃𝑟𝑒𝑓)

𝑅𝑇) =∏(𝑃𝑦𝑖)

𝜈𝑖

𝑁𝐶

𝑖=1

(2.15)

3 EXPERIMENTS IN DILUTED PACKED BED REACTORS

3.1 INTRODUCTION

3.1.1 Choice of the packed bed reactor

Catalyst testing can be carried out in many different types of reactors. In general, a series

of criteria should be satisfied for a correct measurement of the intrinsic properties of the

catalyst: sufficient contact between the reactants and the catalyst, absence of heat and

mass transfer limitations, and a well-defined residence time distribution under

isothermal conditions [35].

The fixed bed reactor shows a series of advantages: it is simple to build and use, it

guarantees good fluid-catalyst contact, it is inexpensive and can be used for both gas-

and liquid-phase systems. It requires small amounts of catalyst, and deactivation

phenomena can be detected directly under steady-state operation. Still, concentration

and temperature gradients, both at the reactor and at the particle scale, might affect the

quality of the experiments in presence of low flow rates, and care must be taken in

ensuring plug-flow behaviour. Temperature gradients are usually the most critical

aspect, due to the poor heat conductivity of gas-solid packed beds: diagnostic criteria for

the detection of heat transport limitations will be discussed in 3.2.

3.1.2 Reactors used in this work

Three different reactors were used for the study of CO PrOx in diluted bed systems.

Name Mass of catalyst

[g]

Mass of

diluent [g]

Conditioned in

H2

Reference

GHSV

[NL/h/kg]

BED1 0.10 0.25 NO 80000

BED2 0.10 0.70 YES 240000

BED3 0.05 0.70 YES 160000

Table 3.1: Packed bed reactors used in this Thesis work.

As it can be seen from the table, the three reactors are characterized by an increasing

diluent to catalyst ratio (from 2.5, to values close to 7 and 14, respectively). This is due


to the fact that, as it will be discussed later, oddly high conversions were observed even

at very low temperatures, consistently with an extremely high activity of the catalyst.

Another reason is that, due to the strong exothermicity of the oxidation reactions, the

temperature of the system was difficult to control. This could result in strong

temperature gradients across the catalytic bed, which can be reduced by increasing the

dilution ratio (as it will be discussed in 3.2.3).

Figure 3.1: BED1.

The reactors are all very similar and consist in an approximately 30-cm long quartz tube,

with a 6 cm internal diameter and a 2 mm wall thickness. The tube is characterized by

the presence of a shrinkage in the bottom part, in order to increase the velocity of the

outlet gas stream, and minimize possible homogeneous contributions [36].

The procedure which was followed in the preparation of the reactors is the following: a

layer of quartz wool was inserted at the bottom part of the reactor. Then, on the top of

it, a mixture of catalyst and diluent of known mass was poured inside the tube, while

carefully keeping the thermocouple at an height corresponding to half of the bed length.

Despite SiC being the diluent of choice in a previous Thesis work [30], quartz was

selected as the inert for these experiments. The reason for this is the possibility of SiC

having a catalytic activity itself with respect to oxygen, thus affecting the experimental

results. Another layer of quartz wool separates the catalytic bed from a layer of

irregularly-shaped quartz crystals, with the function of homogenizing the inlet flow, and

reducing the risk of by-pass [36]. A final quartz wool layer is inserted on the top.

Before the experiments with the catalyst, a blank test (empty tube) was performed in

order to check for the presence of any homogeneous reaction: however, this was not the

case.

INTRODUCTION 51

3.1.3 Operating conditions

For the three beds, a reference composition of 40% hydrogen, 1% CO, 1% O2 and

complementary nitrogen was selected. The amounts of CO and O2 reflect the ones which

can be expected in a real CO PrOx reactor for the application in fuel cell systems: despite

being 0.5% O2 the stoichiometric concentration, oxygen is usually fed to the preferential

oxidation stage in sovrastoichiometric amounts, in order to guarantee a thorough

consumption of CO. The amount of hydrogen was selected referring to the outlet

composition of experiments previously conduted on the water gas shift reaction [37]. In

the literature, experiments have been carried out at many different possible H2

concentrations. In any case, one should expect a large excess of hydrogen to be present

with respect to the other reactants (typical H2/CO ratios exceed 100:1 at the outlet of the

water gas shift stage and can increase up to 50000:1 as the PrOx reaction proceeds to

completion [5]).

Initially, 40 °C-300 °C was the temperature range of choice, since CO PrOx reactors are

usually operated at low temperatures. However, even from the very first experiments, it

was clear that a reliable investigation of the 40 °C-100°C temperature range was not

feasible. This could be observed not only from the poor reproducibility of data at low

temperatures, but also from the sudden temperature increase at the beginning of every

experiment. The set point of the oven was set equal to 40 °C before opening the reactor:

then, as the reactant mixture was sent on the bed, the temperature of the catalyst

increased up to a value which depended on the experimental conditions.

This made it necessary to follow a specific procedure: in order to get to the desired

temperature, the set point of the oven was increased 5 °C at a time, every twenty seconds,

from 0 °C up to the first temperature, in order to avoid abrupt temperature increases. It

should be noted that the lines inside which the reactants were fed to the reactor were

heated up to 120 °C, a temperature which is much higher than the set point of the oven:

however, keeping the heating tapes off and turning them on only at the beginning of the

experiment did not change the situation. Hence, the temperature increase is probably

related to the strong exothermicity of both the CO and the H2 oxidation reactions and to

the poor capability of the packed bed of efficiently dissipate the reaction heat. This was


one of the reasons which eventually made it necessary to move to an annular

configuration (see Chapter 4).

3.1.4 Apparent deactivation of the catalyst

Another interesting phenomenon which could be observed from the very beginning is

the apparent deactivation of the catalyst, resulting in a decrease in the conversion of both

CO and O2 with the on-stream time. As explained in 2.2.2, under steady-state conditions,

the flow rate of each species is estimated by averaging the results of three different

injections, whose results do not vary significantly, or show any particular trend. In this

study, however, the conversions systematically decreased from one injection to another,

especially at the lowest temperatures.

A possible explaination for this temporal deactivation might be again the poor heat

transfer ability of the packed bed system: the temperature indicated by the

thermocouple, located at half of the bed length, did not reflect the actual temperature on

the surface of the catalyst, representative of the reaction rate and progressively

decreasing as heat was dissipated from the bed to the oven.

Figure 3.2.: Conversion drift at 60 °C over a 150-minute time period (BED3). GHSV=160000 NL/h/kg. Inlet

composition: 40% H2, 1% CO, 1% O2.

However, this cannot be the only explaination. As it will be discussed in 4.2.1, the same

phenomenon would have later been observed in the annular system, which is on the

contrary characterized by an efficient heat dissipation. Hence, even if the exothermicity

might play a role in this, it is more probable that this conversion decrease is actually the

20

25

30

35

40

45

50

55

0 10 20 30 40 50

CO

co

nve

rsio

n[%

]

# injection

T = 60 °C

DIAGNOSTIC CRITERIA FOR HEAT TRANSPORT LIMITATIONS 53

visible effect of very slow stabilization dynamics, associated to an extremely low reaction

rate. The phenomenon is also cited in some works [5], [22], [38].

Some authors [39] state that the significant deactivation of the catalyst with the time of

stream is not related to a modification in the intrinsic activity of the catalyst (since the

pseudo-first-order kinetic contant of the reaction is more or less unchanging), but to a

decrease in the surface concentration of CO2 intermediates, related to the available active

Pt sites. The initial activity of the catalyst could be completely restored by treating the

catalyst in a hydrogen stream.

Other authors [18] confirmed that this apparent deactivation is perfectly reversible:

indeed, the initial reaction rates and selectivity were restored simply by treating the

sample in a flow of He at 623 K for 900 s. As a possible explaination for the phenomenon,

the authors assume the progressive increase in chemisorbed CO coverages up to the

steady-state value, in the presence of other species: the kinetic consequences can be

considered equivalent to a gradual increase in the adsorption constant of CO.

In later experiments in the diluted packed bed system, injections were performed for up

to one hour at the first temperature in order to obtain at least a partial stabilization of the

system. This effect was investigated in more depth in BED3 (3.3.3).

3.2 DIAGNOSTIC CRITERIA FOR HEAT TRANSPORT LIMITATIONS

3.2.1 Introduction

In principle, kinetic data should be collected in fixed beds operated as integral reactors

under rigorous isothermal conditions. However, the presence of temperature gradients

can cause a significant deviation from the ideal isothermal, plug-flow behaviour in fixed

bed reactors [40], thus altering the experimental results. Thermal gradients might be

intraparticle, thus related to the heat transfer inside the catalytic pellet; interphase,

between the bulk of the gas phase and the surface of the solid, possibly at a much higher

temperature due the presence of an exothermic reaction; interparticle, between the fluid

and the wall.

A criterion to prove the absence of intraparticle temperature gradients was developed

by Anderson, by imposing a <5% deviation of the observed rate from the isothermal one.


However, Fulton and Crosser proved that for gas-solid systems the intraparticle heat

transfer resistance is negligible with respect to the interphase one. Thus, the particle itself

can be usually deemed isothermal. Diagnostic criteria similar to the one of Anderson can

be derived for the interphase, and interparticle transport.

3.2.2 Interphase transport

A criterion [40] for the detection of interphase heat transfer limitations can be obtained

by the perturbation approach. By assuming an Arrhenius-type dependence on the

temperature for the reaction rate

ℜ = 𝐴𝑒−𝐸 𝑅𝑇⁄ 𝑓(𝑐) (3.1)

then the reaction rate at a temperature 𝑇 close to 𝑇0 (the temperature of the bulk fluid)

can be expressed through a Taylor expansion around 𝑇0:

ℜ = ℜ0 (1 +𝑇 − 𝑇0𝑇0

𝐸

𝑅𝑇0) (3.2)

where ℜ0 is the reaction rate at 𝑇0. By neglecting heat conduction to adjacent touching

particles, the energy balance for the particle can be written as:

𝑞ℜ4𝜋

3(𝑟𝑝)

3 = ℎ(𝑇 − 𝑇0)4𝜋(𝑟𝑝)2 (3.3)

where 𝑞 is the absolute value of the heat of the reaction, 𝑟𝑝 the particle radius and ℎ the

heat-solid gas transfer coefficient. Combining the two equations gives:

ℜ

ℜ0= 1 +

𝑞ℜ𝑟𝑝𝐸

3ℎ𝑇02𝑅

(3.4)

The observed rate ℜ, calculated from experimental data, must not deviate from the one

calculated at the bulk gas temperature 𝑇0 by more than 5%, in order for the system to be

deemed isothermal. The resulting criterion to be respected is the following, valid no

matter the presence of intraparticle diffusional limitations:

𝑞ℜ𝑟𝑝

ℎ𝑇0< 0.15

𝑅𝑇0𝐸

(3.5)

Interphase transfer limitations can thus be expected in the case of high heats of reaction,

or for systems operating at low Reynolds numbers (and thus low ℎ).


3.2.3 Interparticle transport

A rigorous derivation of axial and radial temperature profiles would involve the

resolution of partial differential equations: numerical solutions yield approximately

parabolic radial temperature profiles, while the temperature along the axis increases up

to a maximum at the hot-spot and then decreases gradually until reaching again the wall

temperature 𝑇𝑤.

A criterion for the detection of a radial heat transport limitation was also proposed in

[40]. By referring to a bed with a sufficiently high length to particle diameter ratio, axial

conduction can be neglected. Hence, the energy balance at the hotspot can be written as:

𝑘𝑒 (𝑑2𝑇

𝑑𝑟2+1

𝑟

𝑑𝑇

𝑑𝑟) = (−∆𝐻)

(1 − 휀)

(1 − 𝑏)ℜ = (−∆𝐻)ℜ′ (3.6)

where the sensible term has been neglected. Moreover, the local rate is defined as:

ℜ′ =(1 − 휀)

(1 − 𝑏)ℜ (3.7)

where 휀 is the void fraction and 𝑏 the dilution (inert to catalyst) ratio. By defining the

dimensionless temperature and radius (𝑅0 is the radius of the reactor):

𝜃 =𝐸(𝑇 −𝑇𝑤)

𝑅𝑇𝑤2 (3.8a)

𝑢 =𝑟

𝑅0 (3.8b)

the analytical solution derived for the radial temperature profiles can be expressed as:

𝜃 − 𝜃𝑚𝑎𝑥 = −2ln(𝐵𝑢2 + 1) (3.9a)

𝐵 =𝛿

8exp(|𝜃𝑚𝑎𝑥|) (3.9b)

𝛿 =(−∆𝐻)ℜ′𝑤𝑅0

2𝐸

𝑘𝑒𝑇𝑤2𝑅

(3.9c)

where 𝜃𝑚𝑎𝑥 is the maximum dimensionless temperature at the hotspot.

By imposing the boundary condition 𝑢 = 1 for 𝜃 = 0 (i.e.: the temperature is equal to 𝑇𝑤

at the wall), the following relation is derived:

𝛿 = 8[exp(−0.5|𝜃𝑚𝑎𝑥|) − exp(−|𝜃𝑚𝑎𝑥|)] (3.10)


When the radial temperature gradient is just starting to become significant, 𝛿 is small

and so is 𝐵. Hence, the following approximation can be used:

𝜃 − 𝜃𝑚𝑎𝑥 = −2𝐵𝑢2 (3.11)

Finally, the following expression for 𝜃𝑚𝑎𝑥 can be obtained:

𝜃𝑚𝑎𝑥 =𝛿 4⁄

1 − 𝛿 4⁄≈ 𝛿 4⁄ (3.12)

The heat transfer resistance at the wall is usually not negligible in small laboratory

reactors (𝑅0 𝑟𝑝⁄ < 100), and aggravates the interparticle heat transport. In this case, the

proper boundary condition is given by:

𝑘𝑒𝑑𝑇

𝑑𝑟|𝑟=𝑅0

= ℎ𝑤(𝑇 − 𝑇𝑤)|𝑟=𝑅0 (3.13)

In adimensional terms,

𝑑𝜃

𝑑𝑢|𝑢=1

=ℎ𝑤𝑅0𝑘𝑒

𝜃|𝑢=1 =𝐵𝑖𝑤2

𝑅0𝑟𝑝𝜃|𝑢=1 (3.14)

where 𝐵𝑖𝑤 =ℎ𝑤𝑑𝑝

𝑘𝑒 is the so-called Biot number, an adimensional quantity expressing the

ratio between the heat transfer at the wall and the one at the center of the bed.

Using this boundary condition, the following expression is obtained for 𝜃𝑚𝑎𝑥:

𝜃𝑚𝑎𝑥 =𝛿

4(1 +

4𝑘𝑒ℎ𝑤𝑅0

) =𝛿

4(1 +

8

𝐵𝑖𝑤

𝑟𝑝

𝑅0) (3.15)

which is negative in the case of endothermic reactions.

Finally, the effect of the radial gradients on the reaction rate can be assessed by

substituting eqs. 3.8a and 3.11 into equation 3.2, and integrating across the cross section:

𝜋(1)2ℜ′̅̅ ̅ = ℜ′𝑤∫ (1 + 𝜃𝑚𝑎𝑥 − 2𝐵𝑢2

1

0

)2𝜋𝑢𝑑𝑢 (3.16)

where ℜ′̅̅ ̅ is the average reaction rate at the cross section of the hot-spot. It follows that:

ℜ′̅̅ ̅

ℜ′𝑤= 1 + 𝜃𝑚𝑎𝑥 − 𝐵 (3.17)

The average rate ℜ′̅̅ ̅ should not deviate from ℜ′𝑤 (the local rate at 𝑇𝑤) by more than 5%,

in order for the system to be deemed isothermal. Hence, the following relationship must

hold:


𝑞ℜ′̅̅ ̅𝑅0

2

𝑘𝑒𝑇𝑤<

0.4𝑅𝑇𝑤 𝐸⁄

(1 + 8 𝑟𝑝 𝑅0𝐵𝑖𝑤⁄ ) (3.18)

which simplifies to

𝑞ℜ′̅̅ ̅𝑅0

2

𝑘𝑒𝑇𝑤< 0.4

𝑅𝑇𝑤𝐸

(3.19)

if the heat transfer resistance at the wall is negligible.

Reducing the reactor diameter (present at the numerator in the left hand side in the

interparticle criterion, but also affecting the gas velocity and thus the transport

properties) and the particle radius might help in the minimization of heat transport

limitations. Diluting the bed is also beneficial towards the reduction of radial

temperature gradients, since it results in a smaller rate per unit volume: however,

dilution might significantly decrease the gas velocity and be the source of undesired

bypass phenomena.

3.2.4 Estimation of the transport properties

𝑘𝑒 is the effective thermal conductivity across the bed. It is usually expressed as the sum

of two terms: a static contribution and a dynamic one, related to the gas-solid interaction:

𝑘𝑒𝑘𝑔=𝑘𝑠𝑡𝑎𝑡𝑘𝑔

+𝑘𝑐𝑜𝑛𝑣𝑘𝑔

(3.20)

where 𝑘𝑔 is the thermal conductivity of the gas mixture. For 𝑅𝑒 > 40, the static

contribution can be calculated as (𝑘𝑝 is the thermal conductivity of the particle) [41]:

𝑘𝑠𝑡𝑎𝑡𝑘𝑔

= 휀𝑏 +1 − 휀𝑏

0.220휀𝑏2 + 2/3

𝑘𝑔𝑘𝑝

(3.21)

The dynamic contribution is usually defined as:

𝑘𝑐𝑜𝑛𝑣𝑘𝑔

=1

𝑃𝑒𝑟𝑖𝑓𝑅𝑒𝑝𝑃𝑟 (3.22)

where

𝑅𝑒𝑝 =𝜌𝑔𝑣𝑑𝑝

𝜇𝑔 (3.23)

𝑃𝑟 =𝜇𝑔𝑐𝑝,𝑔

𝑘𝑔 (3.24)


𝑃𝑒𝑟𝑖𝑓 = 8.65 [1 + 19.4 (𝑑𝑝

𝑑𝑡)] (3.25)

The density of the gaseous mixture is calculated through the ideal gas law and 𝑐𝑝,𝑔 as

∑𝑥𝑖𝑐𝑝,𝑖, while both the viscosity and the conductivity of the gaseous mixture have been

calculated by using the ASALI [42] code.

The heat transfer coefficient at the wall is also the sum of two contributions:

ℎ𝑤 = ℎ𝑤,𝑠𝑡𝑎𝑡 + ℎ𝑤,𝑐𝑜𝑛𝑣 (3.26)

where ℎ𝑤,𝑠𝑡𝑎𝑡 can be calculated as [41]:

ℎ𝑤,𝑠𝑡𝑎𝑡𝑑𝑝

𝑘𝑔= 2휀𝑏 +

1 − 휀𝑏

0.0024 (𝑑𝑡𝑑𝑝)1.58

+ 1/3𝑘𝑔𝑘𝑝

(3.27)

and the dynamic contribution as:

{

ℎ𝑤,𝑐𝑜𝑛𝑣 = 0.0835

𝑘𝑔

𝑑𝑝𝑅𝑒𝑝

0.91 (𝑃𝑟

𝑃𝑟𝑎𝑖𝑟,80°𝐶)

1/3

for𝑅𝑒𝑝 < 1200

ℎ𝑤,𝑐𝑜𝑛𝑣 = 1.23𝑘𝑔

𝑑𝑝𝑅𝑒𝑝

0.53 (𝑃𝑟

𝑃𝑟𝑎𝑖𝑟,80°𝐶)

1/3

for𝑅𝑒𝑝 ≥ 1200

(3.28)

REACTOR HISTORY 59

3.3 REACTOR HISTORY

3.3.1 BED1

BED1 was the first diluted packed reactor used for this Thesis work, and it was used to

gain confidence with the system. A diluent to catalyst ratio of 2.5 was selected, by

referring to a previous Thesis work [30]. Three experiments were carried out inside this

reactor, at two different space velocities (80000 and 200000 NL/h/kg, respectively) in

order to roughly investigate the effect of this process parameter. The results are reported

in Figure 3.3.

The first thing which can be observed from the graph of CO conversion is the irregular

trend at the low temperatures, as discussed in 3.1.3. Apart from that, the results of the

first two tests, which were performed under the same conditions, seem to agree fairly

well above 140 °C. The conversion curve at the lower GHSV is shifted leftwards, as it

would be expected, since a lower GHSV is associated to a larger residence time of the

reactants.

At low temperatures, CO conversion is almost constant (but very far from zero, contrary

to what can be found in most literature works) for a wide temperature range. Then, it

increases monotonically up to a certain temperature, reaching a maximum (which is

unitary conversion for both a GHSV of 80000 NL/h/kg and 200000 NL/h/kg).

It is worth noticing that the conversion curves go over the equilibrium curve. While this

might appear counterintuitive at first, it can actually be explained by the occurrence of

the non-equilibrated oxidation reactions. The decrease in the conversion of CO after a

maximum has been reached is due to the reverse water gas shift reaction

H2 + CO2 → H2O + CO

which takes place only when all the oxygen has been consumed.

Contrary to CO conversion, oxygen conversion monotonically increases up to the total

conversion, which is reached at a lower temperature (around 200 °C) in the case of the

lower space velocity, as expected.


50 100 150 200 250 300

0

20

40

60

80

100

Equilibrium

GHSV = 80000 NL/h/kg (#1)

GHSV = 80000 NL/h/kg (#2)

GHSV = 200000 NL/h/kg

CO

co

nve

rsio

n (

%)

Temperature (°C)

50 100 150 200 250 300

0

20

40

60

80

100

Equilibrium

GHSV = 80000 NL/h/kg (#1)

GHSV = 80000 NL/h/kg (#2)


O2 c

on

ve

rsio

n (

%)

Temperature (°C)

50 100 150 200 250 300

0

20

40

60

80

100

Equilibrium

GHSV = 80000 NL/h/kg (#1)

GHSV = 80000 NL/h/kg (#2)


Se

lectivity to

CO

2 (

%)

Temperature (°C)

0 100 200 300

0,000

0,002

0,004

0,006

0,008

0,010

0,012

Equilibrium

GHSV = 80000 NL/h/kg (#1)

GHSV = 80000 NL/h/kg (#2)


yC

O2

Temperature (°C)

0 100 200 300

0,000

0,002

0,004

0,006

0,008

0,010

0,012

yH

2O

Temperature (°C)

Figure 3.3: Effect of the GHSV in BED1. Inlet composition: 40% H2, 1% CO, 1% O2.

REACTOR HISTORY 61

Oxygen consumption alone is not a very indicative parameter, since it provides no

information about how it is being consumed. Hence, it is worth focusing on the graph of

the selectivity of oxygen to CO. Selectivity is an indicator of how much oxygen is

consumed in the CO oxidation on the basis of the total amount of consumed oxygen. As

previously stated, it is calculated as:

𝑆𝐶𝑂2 =0.5(�̇�𝐶𝑂,𝑖𝑛 − �̇�𝐶𝑂)

�̇�𝑂2,𝑖𝑛 − �̇�𝑂2

Thus, a selectivity lower than the unity indicates that oxygen is being consumed to

produce a species other than carbon dioxide, in this case water. Here, selectivity towards

CO oxidation is monotonically decreasing as the temperature increases, even if the

decrease is significant up to about 150 °C, and weaker at higher temperatures. The

selectivity at low temperatures appears to be significantly higher for a GHSV=200000

NL/h/kg. At higher temperatures, the selectivity remains more or less constant at a value

between 50% and 60%, independently of the space velocity.

yCO2 has the same trend as the conversion of CO. The amount of CO2 is slightly higher

than the one of water until a maximum in the conversion of CO is reached. Surprisingly

enough, CO oxidation seems indeed unaffected by the presence of hydrogen, which is

also competing for oxygen and whose oxidation reaction is known to proceed very fast

towards total conversion even at low temperatures. An explaination for this could be

that CO is saturating the surface of the catalyst, hindering hydrogen oxidation. As the

temperature becomes higher and CO desorption becomes more and more favoured, an

increasing portion of oxygen is consumed in the oxidation of hydrogen.

Contrary to CO2, the amount of water monotonically increases with the temperature.

The yH2O plot reflects the trend of the selectivity, since the presence of water is the direct

indicator of the occurrence of hydrogen oxidation, and thus of the consumption of

oxygen in a reaction other than CO oxidation. It would be very hard to derive any piece

of information starting directly from hydrogen conversion, due to the large experimental

error associated to this parameter. In the case of the lower space velocity, water is

produced even at 60 °C. On the contrary, at the higher space velocity it starts being

produced in meaningful amounts only above 120 °C, meaning that the reaction is still

too slow to be observed at the lower temperatures.


3.3.2 BED2

A deeper investigation on the effect of the GHSV and a study of the effect of oxygen

concentration on the kinetics was carried out in BED2. This reactor is characterized by

the same mass of catalyst as BED1, but a diluent to catalyst ratio (7:1) which is almost

three times and a bed length which is twice (2 cm) the one of the previous bed. The strong

exothermicity of the oxidation reactions was diffucult to deal with inside BED1. By

increasing the mass of inert, the catalyst density in terms of mass of catalyst per unit of

volume of the reactor is reduced, and so are the reaction rate and the rate of heat

production.

The reactor was subjected to the following conditioning procedure in a 40% H2, 60% N2

stream: the temperature was first increased from 30 °C up to 380 °C, at a rate of 3 °C/min.

Then, it was kept at 380 °C for 4 hours and lastly it was cooled down in the reducing

stream up to ambient temperature. Actually, as reported on the product sheet, the

catalyst is said to require no conditioning: however, this was carried out anyway, in

order to check whether it could somehow impact the activity of the catalyst.

Before proceding with the investigation of the effect of GHSV and yO2, a couple of tests

performed under the same operating conditions were carried out in order to verify the

stability of the system: the results agree fairly well. Again, the conversion showed the

previously described decreasing trend injection after injection, going from an initial

value of 23% to a roughly stable value below 10% in 115 minutes (Figure 3.4).

Figure 3.4: Conversion drift in BED2. GHSV=240000 NL/h/kg. Inlet composition: 40% H2, 1% CO, 1% O2.

Injections from 6 up to 15 were carried out around 30 mins after the first five ones, at 10-minute intervals.

0,00

5,00

10,00

15,00

20,00

25,00

0 2 4 6 8 10 12 14 16

X_C

O [

%]

# INJECTION

REACTOR HISTORY 63

0 50 100 150 200 250 300 350

0

20

40

60

80

100




CO

co

nve

rsio

n (

%)

Temperature (°C)

50 100 150 200 250 300

0

20

40

60

80

100




O2 c

on

ve

rsio

n (

%)

Temperature (°C)

0 50 100 150 200 250 300 350

0

20

40

60

80

100




Se

lectivity to

CO

2 (

%)

Temperature (°C)

Figure 3.5: Effect of the GHSV in BED2. Inlet composition: 40% H2, 1% CO, 1% O2.

With regard to the effect of the space velocity, the results reflect the ones obtained for

BED1. They are reported in Figure 3.5. Again, CO conversion increases up to unitary

conversion (the lower the space velocity, the lower the temperature at which the

maximum is reached), going over the equilibrium curve, while oxygen conversion

increases monotonically, with a more abrupt increase in the case of the high GHSV and

a smoother trend in the case of the low GHSV.

The selectivity clearly seems to be independent of the space velocity. In particular, it

approaches 100% at low temperature, an indication of the fact that CO oxidation starts

before H2 oxidation. The selectivity rapidly decreases at the onset of hydrogen oxidation:

by 150 °C, as the equilibrium is approached, it is around 50%. A further temperature

increase leads to a slight decrease in the selectivity, again in line with the thermodynamic

equilibrium.


A study on the effect of oxygen concentration on the kinetics of CO preferential oxidation

was also carried out using BED2. By keeping constant the concentration of CO (1%) and

the GHSV (240000 NL/h/kg), three different concentrations of oxygen were analyzed:

0.5%, 0.75% and 1%.

An experiment with a 2% concentration was also started: however, as soon as the

reactants entered the reactor, the temperature started increasing very rapidly from the

set point value (40 °C) up to 280 °C in less than one minute, thus forcing to stop the

experiment. The onset of significant temperature gradients is presumably responsible

for the uneven heating of the catalytic bed: thus, only qualitative, non-rigorous

considerations can be made on this set of experiments.

As it can be observed from the graphs in Figure 3.6, the conversion of CO is higher in

the presence of a higher concentration of oxygen at the same temperature. While a

maximum can be clearly seen in all the three curves, total conversion of CO is not

attained for a 0.5% and 0.75% oxygen content. The conversion of oxygen, on the contrary,

is again monotonically increasing with the temperature and seems not to depend on the

concentration of this species: after a modest increase in the conversion up to 120 °C, a

sharp rise in the curve can be observed above this temperature, with an inflection point

around 160 °C and again a change in the slope after 220 °C, when total conversion of

oxygen is reached. The trend of the selectivity is not very clear, indicating a possible

increase of the selectivity as the concentration of oxygen decreases.

Interestingly, the concentration of water seems to be significantly higher in the case of

1% O2 concentration. Again, yH2O slightly increases up to 200 °C, where a change in the

slope of the curves can be observed. This increase in the amount of water is associated

to the occurrence of hydrogen oxidation and is concomitant to an almost complete

consumption of oxygen.

The trend of yCO2 reflects the one of CO conversion: the concentration of carbon dioxide

decreases as the concentration of oxygen decreases, and so does the one of water. This

indicates that both the oxidation rates depend on the concentration of oxygen with a

positive order.

REACTOR HISTORY 65

0 50 100 150 200 250 300 350

0

20

40

60

80

100 CO 1% - O2 0.5%

CO 1% - O2 0.75%

CO 1% - O2 1%

CO

co

nve

rsio

n (

%)

Temperature (°C)

50 100 150 200 250 300

0

20

40

60

80

100 CO 1% - O2 0.5%

CO 1% - O2 0.75%

CO 1% - O2 1%

O2 c

on

ve

rsio

n (

%)

Temperature (°C)

0 50 100 150 200 250 300 350

0

20

40

60

80

100

Se

lectivity to

CO

2 (

%)

Temperature (°C)

0 100 200 300

0,000

0,002

0,004

0,006

0,008

0,010

0,012

yC

O2

Temperature (°C)

0 100 200 300

0,000

0,002

0,004

0,006

0,008

0,010

0,012

yH

2O

Temperature (°C)

Figure 3.6: Effect of oxygen concentration in BED2. GHSV=240000 NL/h/kg.


The strong exothermicity of the reaction rose some doubts about the possible presence

of any interphase or interparticle temperature gradient. Hence, before performing

further experiments on BED2, Mears’ criteria on interphase heat transport limitations:

𝑞ℜ𝑟𝑝

ℎ𝑇0< 0.15

𝑅𝑇0𝐸

and on radial heat transport limitations:

𝑞ℜ′̅̅ ̅𝑅02

𝑘𝑒𝑇𝑤< 0.4

𝑅𝑇𝑤𝐸

were applied in order to check whether any temperature gradient was affecting the

experiments.

As it can be observed from Figure 3.7, radial temperature gradients were present inside

the reactor, according to Mears’ criterion on interparticle heat transport limitations.

Hence, BED2 was not used any further and a new bed was produced.

Figure 3.7: Check on the presence of interphase and radial temperature gradients in BED2 (GHSV=240000 NL/h/kg,

40% H2, 1% CO, 1% O2). In blue: the term at the left hand side of each criterion (see 3.2.2 and 3.2.3). In red: the

threshold.

REACTOR HISTORY 67

3.3.3 BED3

The last attempt of producing a diluted packed bed reactor is BED3. This reactor is

characterized by a diluent to catalyst ratio (14:1) which is twice the one of the previous

reactor. BED3 was subjected to the same conditioning procedure as BED2, this time in

50% hydrogen. After that, BED3 was subjected to three experiments, carried out under

the same reaction conditions (H2:CO:O2=40:1:1, GHSV=160000 NL/h/kg) and named

reference tests, aimed at checking the stability of the system and the reproducibility of

the data. As it can be seen in Figure 3.8, even if the curves are similar and reflect the

trend observed for CO conversion also in the previous beds, there are significant

differences between them, especially at low temperatures.

0 50 100 150 200 250 300 350

0

20

40

60

80

100

Reference test #1

Reference test #2

Reference test #3

CO

convers

ion (

%)

Temperature (°C)

Figure 3.8: Trend of CO conversion for the three reference tests carried out in BED3. Inlet composition: 40% H2, 1%

CO, 1% O2. GHSV=160000 NL/h/kg.

BED3 was also exploited for the investigation of the extremely slow stabilization of the

conversion. In order to do so, the usual reactant mixture was sent to the reactor.

However, instead of increasing the temperature after some injections and going on with

the experiment, injections were taken until the conversion seemed to be stable. This

procedure was repeated for two different temperatures a day, for three days. Each

stabilization could require up to sixty three-minute-long injections (three hours). The

reacting mixture was still sent into the reactor as the temperature rose from the first to

the second one.

The results are reported in Figure 3.9: the number of injections on the x axis is

representative for the time on stream. Each injection takes three minutes. Between the

first and the second temperature, however, a larger amount of time (around 30 minutes)

is required for the stabilization of the temperature.


Figure 3.9: Conversion drift in BED3 (inlet composition: 40% H2, 1% CO, 1% O2. GHSV=160000 NL/h/kg).

0

20

40

60

80

100

0 2 0 4 0 6 0 8 0 1 0 0

X_C

O [

%]

# INJECTION

DAY ONE

60 °C

90 °C

0

20

40

60

80

100

0 2 0 4 0 6 0 8 0 1 0 0

X_C

O [

%]

# INJECTION

DAY TWO

T=120 °C

150 °C

0

20

40

60

80

100

0 20 40 60 80 100

X_C

O [

%]

# INJECTION

DAY THREE

T=180 °C

210 °C

REACTOR HISTORY 69

As it can be observed, the conversion decreases slowly, in a power-law fashion. The

conversion drift is not only observed at the low temperatures, but at the higher ones, as

well: moreover, the relative decrease seems independent of the temperature.

Consecutive tests carried out on the same day also showed that the partial stabilization

at the second temperature still required a long time, but was associated to a smaller

decrease in relative terms.

The reversible deactivation of the catalyst seems thus caused by the interaction between

the “empty” catalyst surface and the reacting mixture. The phenomenon is compatible

with a self-poisoning effect of CO adsorption on CO oxidation: the initial higher activity

is associated to a weaker inhibition, and a larger number of free active sites. Reaction

rates would then progressively decrease at increasing coverage of CO.

To conclude, a comparison among the data collected in the three beds was carried out

(see Figure 3.10). The performance of the unconditioned BED1 and the one of conditioned

BED2 were compared at 80000 NL/h/kg (BED3 was never operated at such a low space

velocity, since it would have required too small openings of the mass flow controllers):

by neglecting the unreliable points at lower temperatures, no clear difference can be

observed, leading to the conclusion that indeed no conditioning is required by the

catalyst. The performance of BED2 was also compared to the one of BED3 at

GHSV=160000 NL/h/kg, showing good agreement between the two curves.

In order to verify whether the increased dilution ratio could overcome the issues related

to the control of the temperature, Mears’ diagnostic criteria were applied to verify the

presence of interphase and radial gradients in BED3: the results are reported in Figure

3.11. The graph shows again the violation of Mears’ criterion on radial temperature

gradients, even if these were greatly reduced with respect to the previous reactor. No

more experiments were performed on BED3.


50 100 150 200 250 300

0

20

40

60

80

100

BED_1 (unconditioned) - GHSV = 8e4

BED_2 (conditioned) - GHSV = 8e4

CO

co

nve

rsio

n (

%)

Temperature (°C)

0 50 100 150 200 250 300 350

0

20

40

60

80

100

BED_2 - GHSV = 1.6e5

BED_3 - GHSV = 1.6e5

CO

co

nve

rsio

n (

%)

Temperature (°C)

Figure 3.10: Comparison among BED1, BED2 and BED3. Left: comparison between BED1 (unconditioned) and BED2

(conditioned in hydrogen) at GHSV=80000 NL/h/kg. Right: comparison between BED2 and BED3 at GHSV=160000

NL/h/kg. Inlet composition: 40% H2, 1% CO, 1% O2.

Figure 3.11: Check on the presence of interphase and radial temperature gradients in BED3 (GHSV=160000 NL/h/kg,

40% H2, 1% CO, 1% O2). In blue: the term at the left hand side of each criterion (see 3.2.2 and 3.2.3). In red: the

threshold.

REACTOR HISTORY 71

4 EXPERIMENTS IN THE ANNULAR REACTOR

4.1 INTRODUCTION

4.1.1 The annular reactor

For non-equilibrium limited reactions such as CO oxidation, operating at high space

velocities can significantly widen the operating window for intermediate conversions,

which are the relevant ones from a kinetic point of view. The maximum space velocity

which can be reached inside fixed bed reactors, commonly used as laboratory reactors,

is however limited due to the onset of large pressure drops. Moreover, as it was observed

in Chapter 3, temperature gradients due to ineffective heat dispersion can strongly affect

the quality of the results inside such a system [36].

An alternative to packed bed reactors is the annular reactor, a structured reactor

operating under laminar flow conditions. It consists of an alumina tube, onto which an

extremely small amount of catalyst is deposited (the dip coating procedure has been

described in 2.3.4). The gas stream contacts longitudinally the catalytic surface, in

analogy with monolithic reactors [43].

Given the small amount of catalyst, values of GHSVs in orders of magnitude higher than

the ones which can be obtained in a packed bed reactor can be realized: no pressure

drops are present, for the system is operated under laminar regime, and the tortuosity is

null. Moreover, thanks to the presence of additional routes of heat dispersion (such as

radiation), temperature gradients along the axis are strongly reduced.

Besides, differently from the packed bed reactor whose temperature could be monitored

only by means of a fixed thermocouple, the annular reactor allows for the measurement

of the axial temperature profile, making it possible to verify whether the reactor is

actually operating under isothermal conditions. Indeed, the alumina tube can also be

exploited as a thermocouple well for the measurement of axial temperature profiles (see

2.2.3).

The main issue associated with annular reactors is the high void fraction, hence the

possible influence of homogeneous reactions on the results: since blank tests were

INTRODUCTION 73

already performed in the 100 °C-300 °C temperature range and almost no conversion

was observed, this contribution can be neglected in the case of CO and H2 oxidations.

4.1.2 V1

All the experiments carried out inside the annular reactor were performed on the

washcoated alumina tube V1, prepared as described in 2.3.4 and whose properties can

be found in Table 2.4. The tests were specifically aimed at investigating the effect of more

process parameters on the kinetics of the preferential oxidation of CO: a classification

can be found in Table 4.1.

Investigated parameter Tests no.

Effect of the GHSV 20-24

Effect of yCO 26-29

Effect of yO2 31-33, 36, 37

Effect of the GHSV in the absence of H2 41, 42

Effect of yCO in the absence of H2 47, 48

Effect of of yO2 in the absence of H2 40, 44-46

Table 4.1: Tests carried out on V1.

Reference tests performed at GHSV=500000 NL/h/kg, yCO=0.01, yO2=0.01, yH2=0.40

were carried out inbetween the experimental campaigns in order to check for any

possible deactivation of the catalyst, showing good agreement between the datasets and

thus excluding any significant deactivation phenomenon. The results are reported in

Figure 4.1.

100 150 200 250 300

0

20

40

60

80

100

Reference test #1

Reference test #2

Reference test #3

Reference test #4

CO

co

nve

rsio

n (

%)

Temperature (°C)

Figure 4.1: Reference tests performed on V1. Inlet composition: 40% H2, 1% CO, 1% O2. GHSV=500000 NL/h/kg.


4.2 EXPERIMENTS CARRIED OUT IN THE PRESENCE OF HYDROGEN

4.2.1 Stabilization phenomena

The choice of the annular reactor was dictated by the need of obtaining more rigorous

results with respect to the diluted packed bed reactor, not only in terms of minimizing

the temperature gradients, but also in terms of limiting any stabilization phenomena.

Still, also in the case of the annular reactor, the conversion of both CO and oxygen

seemed to decrease injection after injection at constant temperature.

Since the axial temperature profiles were always rather smooth and close to the nominal

value, the hypothesis of a bed temperature very far from the set point of the oven, and

thus of a higher conversion related to a temperature higher than the nominal one, had

to be discarded. The decrease in the conversion is thus most probably related, as already

explained, to the self-poisoning effect of CO, slowly saturating the surface.

0 10 20 30 40 50

0

5

10

15

20

25

30 100 °C

90 °C

80 °C

CO

co

nve

rsio

n (

%)

# injection

Figure 4.2: CO conversion drift at 100 °C, 90 °C and 80 °C on reactor V1. Inlet composition: 40% H2, 1% CO, 1%

O2. GHSV=500000 NL/h/kg.

Again, a test was performed in order to check the extent of this conversion decrease

(Figure 4.2). By neglecting the very first point, what can be observed despite the large

scattering of the data is that CO conversion at 100 °C decreased of more than the 60% in

the time required to perform 50 injections, about 150 minutes, even if this decrease is

much more significant for the first 15 injections. By carrying on with the experiment at

EXPERIMENTS CARRIED OUT IN THE PRESENCE OF HYDROGEN 75

90 °C and 80 °C, a further decrease in the conversion injection after injection could be

observed. The whole test took around 6 hours.

Since the phenomenon could not be limited, being it related to the reaction mechanism

itself, and since obtaining a stable system should have required an amount of time

significantly longer than the time to complete the tests themselves, it was chosen not to

wait for the complete stabilization, and to continue performing the experiments under

unsteady conditions. However, a partial stabilization was obtained at the beginning of

each experiment by waiting at least 30 minutes for measuring the concentration of the

products. By carrying out the same standardized procedure at every experiment, a good

reproducibility of the results was achieved and the three injections taken at each

temperature seemed not to follow any systematic trend.

Figure 4.3: Trend of the outlet flow rate of CO for the three injections taken at each temperature. GHSV=500000

NL/h/kg. Inlet concentration: 40% H2, 1% CO, 1% O2.

Data measured at the lower temperatures were anyway not considered for the

quantitative kinetic study. Indeed, they are those affected by the highest uncertainty,

both due to the fact that they were gathered after a short time on stream, and also since

CO desorption, possibly limiting the self-poisoning effect, is favoured only at higher

temperatures.

0

0,2

0,4

0,6

0,8

1

1,2

0 1 2 3 4

F_C

O_o

ut

[Nm

L/m

in]

# injection

100 °C

120 °C

140 °C

160 °C

180 °C

200 °C

220 °C

240 °C

260 °C

280 °C

300 °C


4.2.2 Effect of the GHSV

Experiments were carried out in order to investigate the effect of the space velocity. Four

GHSVs were used: 300000, 500000, 1000000 and 1500000 NL/h/kg. The composition used

in every test is the reference one: 40% H2, 1% CO, 1% O2, complementary N2. The results

are reported in Figure 4.4.

The trend of the CO conversion is the same which was already seen in the diluted bed

system: the conversion increases up to a maximum, going over the equilibrium curve,

corresponding to a certain temperature, then decreases. For a space velocity of 1000000

NL/h/kg or above, the region of decreasing conversion is not present. Inside this range

of space velocities, total conversion is also never achieved: the maximum conversion

(~97%) is obtained around 260 °C for a GHSV=300000 NL/h/kg. As the space velocity

increases, the curve is shifted rightwards and is characterized by lower conversions, as

expected. Despite the points collected at the lower temperatures showing again a slighly

irregular trend, it can be clearly seen that the conversion measurement taken at 100 °C

is never null, and it is lower at higher space velocities.

The trend of O2 conversion also reflects the one which was already observed inside the

packed bed reactor. Oxygen conversion, differently than CO conversion, increases

monotonically: it is worth noticing that only when unitary O2 conversion is reached (at

260 °C at 300000 NL/h/kg and at 300 °C at 500000 NL/h/kg), a maximum in the CO

conversion curve is present, indicating again the occurrence of rWGS only as oxygen is

depleted, and the enhancement of hydrogen oxidation as CO desorption becomes more

and more favoured with the temperature.

Selectivity clearly shows no dependence on the space velocity at high temperatures. This

can be explained as follows: by assuming plug-flow conditions inside the reactor, the

outlet conversion and selectivity obtained, for example, at 200 and 100 mL/min,

correspond to the ones at ¼ and ½ of the bed length for a flow rate of 50 mL/min [23].

Hence, if the selectivity seems independent on the space velocity, it means that it is more

or less constant along the reactor. Thus, the differential selectivity and the integral

selectivity are the same and so is the ratio between 𝑟𝐻2 and 𝑟𝐶𝑂 along the reactor for a

given temperature. As an additional consideration, the selectivity is clearly higher at

lower temperatures, even its trend with the GHSV below 200 °C is rather unclear.


100 150 200 250 300

0

20

40

60

80

100 GHSV = 3e5 [NL/h/kg]

GHSV = 5e5 [NL/h/kg]


GHSV = 1.5e6 [NL/h/kg]

CO

co

nve

rsio

n (

%)

Temperature (°C)

100 150 200 250 300

0

20

40

60

80

100

O2 c

on

ve

rsio

n (

%)

Temperature (°C)

100 150 200 250 300

0

20

40

60

80

100 GHSV = 3e5 [NL/h/kg]




Se

lectivity to

CO

2 (

%)

Temperature (°C)

100 150 200 250 300

0,000

0,002

0,004

0,006

0,008

0,010

0,012

0,014 GHSV = 3e5 [NL/h/kg]




yC

O2

Temperature (°C)

100 200 300

0,000

0,002

0,004

0,006

0,008

0,010

0,012

0,014 GHSV = 3e5 [NL/h/kg]




yH

2O

Temperature (°C)

Figure 4.4: Effect of the GHSV in V1. Inlet composition: 40% H2, 1% CO, 1% O2.


By observing the axial temperature profiles, two things can be noticed. The first one is

that the axial temperature profiles (reported in Figure 4.5) show a weak hot-spot, related

to the exothermicity of the oxidation reactions, but are basically flat (the maximum

difference from the set point temperature is about 25 °C at 300 °C for a GHSV=1500000

NL/h/kg, with a maximum gap between axial temperatures of 14 °C): thus, the system

can be indeed considered isothermal. Hence, the high conversions observed even at 100

°C seem not related to a bed temperature significantly different from the nominal one

due to the exothermicity of the reaction.

The second thing which can be observed is that the temperature difference between oven

and bed (reported in Figure 4.6) tends to be higher at higher space velocities. This is due

to the fact that the reactant stream is entering the reactor at a temperature (120 °C,

regulated through the heating tapes) which is lower than the nominal one of the oven:

the higher the flow rate, the less relevant the cooling effect associated to the gas flow due

to a lower residence time. This can also be seen from the fact that the hot-spot is more

pronounced at the higher space velocity.

32 33 34 35 360

50

100

150

200

250

300

350

400GHSV = 3e5

T_

be

d [

°C]

axial coordinate [cm]

catalytic bed

100

120

140

160

180

200

220

240

260

280

300

32 33 34 35 360

50

100

150

200

250

300

350

400GHSV = 1.5e6

T_

be

d [

°C]


catalytic bed

100

120

140

160

180

200

220

240

260

280

300

Figure 4.5: Axial temperature profiles for the tests performed at 300000 and 1500000 NL/h/kg.

32 33 34 35 36-15

-10

-5

0

5

10

15

20

25GHSV = 300000 [NL/h/kg]

100

120

140

160

180

200

220

240

260

280

300

T_bed -

T_oven [°C

]


catalytic bed

32 33 34 35 36-15

-10

-5

0

5

10

15

20

25GHSV = 1500000 [NL/h/kg]

100

120

140

160

180

200

220

240

260

280

300

T_

be

d -

T_

ove

n [

°C]


catalytic bed

Figure 4.6: Axial temperature difference between the catalytic bed and the oven for the tests performed at 300000 and

1500000 NL/h/kg.


4.2.3 Effect of yCO

The effect of the concentration of carbon monoxide on the kinetics of CO PrOx was

investigated by performing a series of tests at different CO concentration, ranging from

0.5% up to 4%: the results are reported in Figure 4.7. Oxygen concentration is 1% and the

GHSV is 500000 NL/h/kg for all tests. This study was not carried out inside the diluted

packed bed reactor.

CO conversion clearly increases as its concentration decreases, at least up to 200 °C,

when a change in the trend of the curves at lower concentration of CO can be observed.

In fact, while in the case of a 0.5% and 1% concentration a maximum is present at 220 °C

and 280 °C, respectively, the curves obtained for yCO=4% and yCO=2% are only

monotonically increasing up to 300 °C. Total CO conversion is never achieved: the

maximum conversion (around 97%) is obtained for a CO concentration of 0.5%. The

effect of CO concentration on the conversion is indicative of a reaction order significantly

lower than 1.

Oxygen conversion monotonically increases with the temperature in every test, but

shows a progressive decrease at increasing CO concentration. It is worth noticing that,

for a 2% and a 4% CO concentration, total O2 conversion is not reached at 300 °C.

The amount of CO2, the product of CO oxidation, follows the same trend as CO

conversion. Its production slightly decreases at increasing CO concentration: the effect

is more important at intermediate temperatures (180-200 °C). At higher temperatures,

the presence of the stoichiometric constraint causes a flattening in the trend of the

concentration. The curves have a tendency to overlap, especially at lower temperatures,

despite yCO2 being slightly higher in the case of lower CO concentrations: hence, it can

be concluded that the reaction rate most probably depends on CO concentration with a

negative reaction order, which seems however close to 0.

The amount of water seems insensitive to the concentration of CO, at least above a 1%

CO concentration. However, the amount of water produced for hydrogen oxidation is

remarkably higher for the lowest CO concentration (0.5%), even at the lower

temperatures: a sharp increase in water concentration can be observed between 180 °C

and 200°C, as oxygen is depleted.


The selectivity seems to depend only slightly on the amount of CO, at least for a 4%, 2%

and 1% concentration. However, the selectivity is significantly lower for a 0.5% carbon

monoxide concentration.

These experimental observations seem to confirm the fact that, in the case of CO PrOx

carried out on Pt-type catalysts, the surface is possibly at least up to a certain

temperature saturated by absorbed CO, hindering oxygen adsorption and the

subsequent oxidation of both of the two species competing for it. For a sufficiently low

concentration of CO both oxidation reactions are favoured, since the species is not able

to saturate the surface anymore.


50 100 150 200 250 300 350

0

20

40

60

80

100 CO 4% - O2 1%

CO 2% - O2 1%

CO 1% - O2 1%

CO 0.5% - O2 1%

CO

co

nve

rsio

n (

%)

Temperature (°C)

50 100 150 200 250 300

0

20

40

60

80

100 CO 4% - O2 1%

CO 2% - O2 1%

CO 1% - O2 1%

CO 0.5% - O2 1%

O2 c

on

ve

rsio

n (

%)

Temperature (°C)

50 100 150 200 250 300 350

0

20

40

60

80

100 CO 4% - O2 1%

CO 2% - O2 1%

CO 1% - O2 1%

CO 0.5% - O2 1%

Sele

ctivity to C

O2 (

%)

Temperature (°C)

100 200 300

0,000

0,002

0,004

0,006

0,008

0,010

0,012

0,014

0,016 CO 4% - O2 1%

CO 2% - O2 1%

CO 1% - O2 1%

CO 0.5% - O2 1%

yC

O2

Temperature (°C)

100 200 300

0,000

0,002

0,004

0,006

0,008

0,010

0,012

0,014

0,016

yH

2O

Temperature (°C)

Figure 4.7: Effect of CO concentration on V1. GHSV=500000 NL/h/kg.


4.2.4 Effect of yO2

In order to study the effect of the concentration of oxygen on the kinetics of CO, a series

of test at different O2 amount were carried out. In particular, yO2 was varied between

0.25% and 4% while CO concentration was always equal to 1% and the GHSV to 500000

NL/h/kg. The results are reported in Figure 4.8.

The results are in good agreement with the ones obtained for the diluted packed bed.

CO conversion clearly depends on oxygen concentration. In particular, higher oxygen

concentrations correspond to higher conversions. The conversion is significantly smaller

at lower temperatures, possibly due to the large CO surface coverage.

Oxygen conversion decreases moderately at increasing oxygen concentration below 175

°C, suggesting a reaction order lower than 1 for this temperature range. On the contrary,

above 175 °C, the conversion of O2 increases as its concentration increases, possibly due

to the larger conversion of CO.

The selectivity seems to be more or less constant and decreasing as T increases for a wide

range of O2 concentrations. However, for a 2% and 4% oxygen concentration, the

selectivity is significantly lower, and the concentration of water produced in hydrogen

oxidation is much higher. Hence, a oxygen concentration significantly higher than the

stoichiometric value seems to be detrimental for the selectivity, even at the low

temperatures.

The amount of both CO2 and H2O is higher at higher oxygen concentration. This is an

indication of a positive order kinetics, possibly first order since the concentration of the

products is increasing more or less linearly (apart from the test at 4% O2 concentration)

with oxygen concentration at the lower temperatures.


100 150 200 250 300

0

20

40

60

80

100

O2 4%

O2 2%

O2 1%

O2 0.75%

O2 0.5%

O2 0.25%

CO

co

nve

rsio

n (

%)

Temperature (°C)

100 150 200 250 300

0

20

40

60

80

100 O2 4%

O2 2%

O2 1%

O2 0.75%

O2 0.5%

O2 0.25%

O2 c

on

ve

rsio

n (

%)

Temperature (°C)

50 100 150 200 250 300 350

0

20

40

60

80

100

Se

lectivity to

CO

2 (

%)

Temperature (°C)

50 100 150 200 250 300 350

0,000

0,002

0,004

0,006

0,008

0,010

0,012 O2 4%

O2 2%

O2 1%

O2 0.75%

O2 0.5%

O2 0.25%

yC

O2

Temperature (°C)

50 100 150 200 250 300 350

0,00

0,01

0,02

0,03

0,04

0,05

0,06

0,07

0,08

0,09

0,10

yH

2O

Temperature (°C)

Figure 4.8: Effect of oxygen concentration on V1. GHSV=500000 NL/h/kg.


4.3 EXPERIMENTS CARRIED OUT IN THE ABSENCE OF HYDROGEN

4.3.1 Introduction

The same sets of experiments performed in the presence of hydrogen (effect of the

GHSV, effect of CO concentration, effect of O2 concentration) were repeated in the

absence of hydrogen, in order to verify to which extent the kinetics of CO oxidation is

impacted by the presence of hydrogen.

4.3.2 Effect of the GHSV

The results of the experiments on the effect of the GHSV in the absence of hydrogen are

reported in Figure 4.9.

The main difference which can be immediately observed in the graph for CO conversion

in the case of the experiments carried out in the absence of hydrogen, is that CO is almost

not at all converted up to 200 °C, no matter which the space velocity. The curves are only

monotonically increasing, since no hydrogen is present in the system and thus neither

rWGS or hydrogen oxidation can occur. As it would be expected, the conversion is

always higher for lower space velocities (and thus higher contact times). The trend of the

curves is also less smooth with respect to the ones obtained in the presence of hydrogen,

showing a sharp increase after 180 °C, as CO oxidation finally starts. For instance, by

observing the curve for a 300000 NL/h/kg GHSV, a conversion increase almost 0-100% is

achieved in a 60-degree interval. 50% conversion is reached around 220 °C, 235 °C and

250 °C in the case of 300000, 500000 and 1000000 NL/h/kg, respectively. Equilibrium total

consumption of CO is achieved for all the three space velocities.

By observing the graph for oxygen conversion, what can be instantly seen is that the

maximum conversion is always the equilibrium one, i.e. 50%. It could not be different

from this, since oxygen cannot be consumed in reactions other than CO oxidation. The

trend of the curves is monotonically increasing and, again, conversion is always higher

for lower space velocities. Equilibrium conversion is achieved at lower temperatures in

the case of lower space velocities. No conversion is observed up to 180 °C, as already

observed in the graph for CO conversion.

EXPERIMENTS CARRIED OUT IN THE ABSENCE OF HYDROGEN 85

What has just been decribed is the proof of the fact that hydrogen is strongly altering the

reactivity of CO, especially at low temperatures: some possible explainations for this

hydrogen-induced enhancement of the reactivity have been discussed in 1.4.3.

100 150 200 250 300

0

20

40

60

80

100 GHSV = 1e6

GHSV = 5e5

GHSV = 3e5

CO

co

nve

rsio

n (

%)

Temperature (°C)

50 100 150 200 250 300

0

20

40

60

80

100

O2 c

on

ve

rsio

n (

%)

Temperature (°C)

100 200 300

0,000

0,002

0,004

0,006

0,008

0,010

0,012

yC

O2

Temperature (°C)

Figure 4.9: Effect of the GHSV in the absence of hydrogen on V1 (squares). The results are compared with the ones of

the experiments carried out in the presence of hydrogen (triangles). Inlet composition: 40% H2, 1% CO, 1% O2.


4.3.3 Effect of yO2

The effect of yO2 in the absence of hydrogen was investigated by performing tests at

0.25%, 0.5%, 0.75%, 1% and 2% concentrations. Hence, both under- and over-

stoichiometric oxygen amounts are taken into consideration. The results are reported in

Figure 4.10. For the sake of clarity, the curves obtained for the corresponding

experiments in the presence of hydrogen are not reported in the graphs.

As it can be observed, CO conversion is higher at higher oxygen concentrations. Again,

almost no conversion is observed until a certain temperature, which appears to be

slightly lower for higher oxygen concentrations. Differently from the curves obtained in

the presence of hydrogen, CO conversion monotonically increases and shows no

decreasing branch: again, no reaction other than CO oxidation can take place. In the cases

of over-stoichiometric oxygen, 100% equilibrium conversion is achieved, at a

temperature which is lower at higher oxygen concentrations.

By observing the graph for O2 conversion, almost no consumption of oxygen is seen until

200 °C, differently from the case of the hydrogen-rich system. The curves are almost

overlapped at low temperature, and then split. Oxygen conversion increases

monotonically, until reaching the equilibrium conversion (except for the 0.25% and the

0.5% tests).

Finally, CO2 concentration is clearly higher for higher oxygen amounts. This trend is an

indication of the fact that the reaction depends with a positive order on the partial

pressure of oxygen.


100 150 200 250 300

0

20

40

60

80

100 O2 2%

O2 1%

O2 0.75%

O2 0.5%

O2 0.25%

CO

co

nve

rsio

n (

%)

Temperature (°C)

50 100 150 200 250 300

0

20

40

60

80

100 O2 2%

O2 1%

O2 0.75%

O2 0.5%

O2 0.25%

O2 c

on

ve

rsio

n (

%)

Temperature (°C)

100 200 300

0,000

0,002

0,004

0,006

0,008

0,010

0,012 O2 2%

O2 1%

O2 0.75%

O2 0.5%

O2 0.25%

yC

O2

Temperature (°C)

Figure 4.10: Effect of oxygen concentration on V1 in the absence of hydrogen. GHSV=500000 NL/h/kg.


4.3.4 Effect of yCO

Tests at 1%, 2% and 0.5% CO concentration were performed in order to investigate the

effect of this parameter on the kinetics. The results are reported in Figure 4.11.

By observing the trend of CO conversion in the absence of hydrogen, what can be

observed is that the conversion is again monotonically increasing, until reaching the

unitary, equilibrium value in the presence of a stoichiometric (1% CO) and over-

stoiochiometric (2% CO) amount of oxygen (whose inlet concentration was always set

equal to 1%). No conversion is seen until around 180 °C for 1% and 2% CO, while in the

case of 0.5% CO concentration the conversion is already close to 20% at this temperature.

The graph for O2 conversion shows that this parameter clearly depends on the

concentration of CO. In fact, while up to 160 °C the curves tend to overlap, from 180 °C

they separate: the conversion is higher at lower CO concentrations. A plateau is reached

both for a 1% and a 0.5% CO concentration, corresponding to the equilibrium

conversion. Differently from the tests carried out in the presence of hydrogen, total

conversion is never achieved, since CO is always present in stoichiometric or over-

stoichiometric amount with respect to the CO oxidation reaction.

Finally, the graph for CO2 concentration also shows that its production rate is enhanced

by lower CO concentrations.


100 150 200 250 300

0

20

40

60

80

100 CO 2%

CO 1%

CO 0.5%

CO

co

nve

rsio

n (

%)

Temperature (°C)

50 100 150 200 250 300

0

20

40

60

80

100 CO 2%

CO 1%

CO 0.5%

O2 c

on

ve

rsio

n (

%)

Temperature (°C)

100 200 300

0,000

0,002

0,004

0,006

0,008

0,010

0,012

0,014

0,016

0,018

0,020

0,022

yC

O2

Temperature (°C)

Figure 4.11: Effect of CO concentration on V1 in the absence of hydrogen. GHSV=500000 NL/h/kg.

5 KINETIC STUDY

5.1 INTRODUCTION

5.1.1 Rate of the reaction

By considering a generic reaction:

𝜈𝐴𝐴 + 𝜈𝐵𝐵 → 𝜈𝐶𝐶 + 𝜈𝐷𝐷

The change in the number of moles of a species per unit of time is defined as the rate of

production (or disappearance) of a species. The following equalities hold:

−1

𝜈𝐴

𝑑𝑛𝐴𝑑𝑡

= −1

𝜈𝐵

𝑑𝑛𝐵𝑑𝑡

=1

𝜈𝐶

𝑑𝑛𝐶𝑑𝑡

=1

𝜈𝐷

𝑑𝑛𝐷𝑑𝑡

and each of these quantities can be considered the rate of the reaction [44], usually

expressed on an intensive basis such as the reaction volume or, in the case of

heterogeneous reactions, the mass of catalyst.

The rate of a reaction can be expressed as a function of the temperature, the pressure and

the concentrations of the species which are present in the system:

ℜ = ℜ(𝑇, 𝑃, 𝐜)

By separating the dependence on the temperature from the one on the concentrations:

ℜ(𝑇, 𝒄) = 𝑘(𝑇) ∙ 𝑓(𝒄)

𝑘(𝑇) is a proportionality factor called kinetic constant, usually expressed through

Arrhenius’ law:

𝑘(𝑇) = 𝐴𝑒𝑥𝑝 (−𝐸𝑎𝑅𝑇)

𝐴 is the so-called pre-exponential factor and 𝐸𝑎 is the activation energy. By this law, by

plotting the logarithm of the kinetic constant against 1 𝑇⁄ , a straight line of slope 𝐸𝑎

𝑅⁄ is

obtained.

A modified form of Arrenhius’ law can also be exploited:

𝑘(𝑇) = 𝐴𝑒𝑥𝑝 [−𝐸𝑎𝑅(1

𝑇−

1

𝑇𝑟𝑒𝑓)]

INTRODUCTION 91

𝑓(𝒄) can be expressed both as a function of the concentrations (usually in the case of

liquid-phase reactions) or of the partial pressures (in the case of gas-phase reactions). For

highly non-ideal systems, fugacities might replace these quantities.

The simplest way to express 𝑓(𝒄) is the following:

𝑓(𝒄) =∏𝑐𝑖𝛼𝑖

𝑁𝐶

𝑖=1

or𝑓(𝒄) =∏𝑃𝑖𝛼𝑖

𝑁𝐶

𝑖=1

where 𝛼𝑖 is the reaction order with respect to species i. The sum of the reaction orders is

defined global order of the reaction. The reaction rate will be expressed as:

ℜ(𝑇, 𝒄) = 𝑘(𝑇) ∙∏𝑐𝑖𝛼𝑖

𝑁𝐶

𝑖=1

This kind of expression is usually defined power-law. Only in the case of elementary

reactions, the reaction orders strictly coincide with the stoichiometric coefficients of the

reactants, so that:

ℜ(𝑇, 𝒄) = 𝑘(𝑇) ∙ ∏ 𝑐𝑖𝜈𝑖

𝑁𝑟𝑒𝑎𝑐𝑡𝑎𝑛𝑡𝑠

𝑖=1

If obtained through the fitting of experimental data, as it is common, the reaction orders

have however no particular physical meaning, since the overall reaction is just the result

of a series of elementary steps. Hence, the reaction orders can be positive, null, or even

negative. A limit of this approach is that a power-law expression is usually unreliable

outside of the investigated range of operating conditions.

An alternative approach is the mechanistic one. A series of elementary steps are

identified: then, by means of reasonable assumptions, the rates of the single steps are

reduced to the rate of a global reaction, which will be function of the concentrations of

stable species. The final expression is usually very accurate, but such a method requires

a deep knowledge of the reacting system. Moreover, it is quite often time-expensive and

heavy from a computational point of view.

Two main things have to be derived from a kinetic analysis: the reaction order, both

global or partial with respect to each of the species, and the kinetic constant (in terms of

pre-exponential factor, and activation energy). The kinetic analysis can be carried out

either in the differential, or in the integral regime.

92 KINETIC STUDY

5.1.2 Kinetic analysis in differential regime

If the concentration of the reactants and the temperature do not change significantly

along the reactor, the reaction rate can also be deemed constant. If this is the case, the

behaviour of the reactor is defined differential.

By considering, for instance, the oxidation of CO alone, the reaction rate can be expressed

as a function of the inlet concentrations:

ℜ = 𝑘𝑃𝐶𝑂𝛼 𝑃𝑂2

𝛽= 𝑘𝑃𝐶𝑂,𝑖𝑛

𝛼 𝑃𝑂2,𝑖𝑛𝛽

and it can be directly calculated from the experimental data starting from the molar flow

rate of the product, and the mass of catalyst in the bed:

ℜ =�̇�𝐶𝑂2𝑊𝑐𝑎𝑡

By separating the effects of CO and oxygen (the subscript in is omitted):

ℜ = 𝑘𝐶𝑂𝑃𝐶𝑂𝛼

ℜ = 𝑘𝑂2𝑃𝑂2𝛽

and linearizing:

ln(ℜ) = ln(𝑘𝐶𝑂) + 𝛼ln(𝑃𝐶𝑂)

ln(ℜ) = ln(𝑘𝑂2) + 𝛽ln(𝑃𝑂2)

By means of a simple linear interpolation, the reaction orders 𝛼 and 𝛽 can be estimated.

It must be underlined that the differential approach can be exploited only as long as the

conversion of the limiting reactant is below a certain threshold, which depends on the

sensitivity of the system. For this analysis, it has been assumed equal to 20%.

Above this threshold, the variation in the reaction rate along the reactor cannot be

deemed negligible anymore, and an integral analysis should be carried out instead.

MATHEMATICAL MODEL OF THE ANNULAR REACTOR 93

5.2 MATHEMATICAL MODEL OF THE ANNULAR REACTOR

5.2.1 Introduction

The kinetic parameters were derived through the comparison of the experimental data

with the results obtained from a model for an isothermal plug-flow reactor. Such model

was developed [45] for the simulation of the catalytic partial oxidation of methane on a

Rh/Al2O3 catalyst inside an annular reactor: hence, the kinetic scheme had to be

extensively modified.

The main features of the model are the following:

• one-dimensional PFR, heterogeneous (all the radial gradient are localized inside

an infinitesimal layer by the catalytic suface);

• negligible axial dispersion (high value of the axial Péclet number);

• laminar flow;

• isothermal (temperature gradients are absent both in the axial direction and

between the bed and the gas phase).

The reactor can be considered isothermal as long as the temperature difference along the

bed does not exceed 10 °C. If such assumption is not valid, the temperature profile is

estimated by means of a fifth-order polynomial in the axial coordinate, as follows:

𝑇(𝑧) = 𝑎1𝑧5 + 𝑎2𝑧

4 + 𝑎3𝑧3 + 𝑎4𝑧

2 + 𝑎5𝑧 + 𝑎6

Both interphase and intraphase mass transfer limitations can be accounted for.

5.2.2 Equations of the model

Since the model is heterogeneous, it can be used to estimate the concentration profiles

along the axis for both the gas phase and the surface of the solid. The number of

unknowns of the problem is 2NC, i.e. NC surface concentrations 𝑥𝑖𝑆 and NC

concentrations 𝑥𝑖𝐵 in the bulk of the gas phase, where NC is the number of components.

To solve the problem, 2NC equations have to be written. In particular:

• NC mass balance equations for the gas phase, in adimensional terms:

𝑃𝑒𝑚,𝑖𝑑𝐹𝑖

∗

𝑑𝑧∗= −

4

1 +1𝑅∗

𝑆ℎ𝑙𝑜𝑐,𝑖(𝑥𝑖𝐵 − 𝑥𝑖

𝑆)𝐹𝑡𝑜𝑡

𝐹𝑡𝑜𝑡0

where 𝑆ℎ𝑙𝑜𝑐,𝑖 is the local Sherwood number (see next paragraph), and:

94 KINETIC STUDY

𝑃𝑒𝑚,𝑖 =�̅�𝐷ℎ𝑦𝑑𝑟

𝒟𝑖

𝐷ℎ𝑦𝑑𝑟 = 𝐷𝑜𝑢𝑡 − 𝐷𝑖𝑛

𝑧∗ =𝑧

𝐷ℎ𝑦𝑑𝑟

𝐹𝑖∗ =

𝐹𝑖

𝐹𝑡𝑜𝑡0

𝑅∗ =𝐷𝑖𝑛𝐷𝑜𝑢𝑡

Of the NC ordinary differential equations, three can be replaced by the atomic

balances for the bulk phase on carbon, hydrogen and oxygen. In fact, being the

atomic balances first-order linear, algebraic equations, the replacement leads to

a significant simplification of the model and thus to a reduction of the

computational time.

∑(𝐹𝑖 − 𝐹𝑖0)𝑛𝐶,𝑖 = 0

𝑁𝐶

𝑖=1

∑(𝐹𝑖 − 𝐹𝑖0)𝑛𝐻,𝑖 = 0

𝑁𝐶

𝑖=1

∑(𝐹𝑖 − 𝐹𝑖0)𝑛𝑂,𝑖 = 0

𝑁𝐶

𝑖=1

• NC continuity equations for the catalyst phase:

𝑆ℎ𝑙𝑜𝑐,𝑖(𝑥𝑖𝐵 − 𝑥𝑖

𝑆) =∑𝜈𝑖,𝑗𝛼𝑖ℜ𝑗

𝑁𝑅

𝑗=1

where ℜ𝑗 is the reaction rate of the j-th reaction, and:

𝛼𝑖 =𝐷ℎ𝑦𝑑𝑟𝑊𝑐𝑎𝑡

𝑆𝒟𝑖𝑐𝑡𝑜𝑡

𝑊𝑐𝑎𝑡 is the mass of catalyst and 𝑆 its surface area. 𝑐𝑡𝑜𝑡 is the total concentration in

the gas phase, calculated under the assumption of mixture of ideal gases as 𝑃/𝑅𝑇.

𝒟𝑖 is the molecular diffusivity of species i and is approximated as the binary

diffusivity of species i in nitrogen, calculated according to the Fuller-Schletter-

Giddings correlation.

The 2NC equations can be solved by means of numerical methods.

MATHEMATICAL MODEL OF THE ANNULAR REACTOR 95

5.2.3 Mass transfer resistances

To account for the presence of interphase mass transfer resistances, the following

expression is used for the estimation of the axial profile of the local Sherwood number:

𝑆ℎ𝑙𝑜𝑐,𝑖(𝑧𝑆ℎ,𝑖) = 𝑆ℎ𝑖𝑛𝑓 + 6.874exp(−71.2𝑧𝑆ℎ,𝑖)(1000𝑧𝑆ℎ,𝑖)−0.35

where

𝑧𝑆ℎ,𝑖 =𝑧∗

𝑃𝑒𝑚,𝑖

𝑆ℎ𝑖𝑛𝑓 = 6.6156 − 1.7548𝑅∗

The functional form of the expression for 𝑆ℎ𝑙𝑜𝑐,𝑖(𝑧𝑆ℎ,𝑖) had already been used in the

literature for the interpolation of the exact solutions of the Graetz-Nusselt problem (and

analogous mass transfer problems) in ducts of various geometries. The coefficients have

been derived by adapting the expression to an annular geometry with boundary

conditions of the third-type [43].

The model is also capable of accounting for intraphase mass transfer limitations by

means of a generalized efficiency factor for oxygen (and also for methane in the original

model [45]). Indeed, the oxidation rates are assumed to be limited by oxygen: this effect

is included by multiplying each of the two rates by the efficiency factor.

At least as a first approximation, internal mass transfer resistances were not considered

in the modelling phase. Thus, the derived kinetic expressions already account for the

presence of diffusion limitations inside the washcoat: being its thickness close to the one

of the active phase of the catalytic pellet, the expression for the reaction rate should still

be valid. However, the impact of internal mass transfer limitations on the results is still

to be verified in a second moment.

5.2.4 Reaction rates

The overall rate of production (or consumption) of species i is calculated as:

𝑟𝑖 =∑𝜈𝑖,𝑗ℜ𝑗

𝑁𝑅

𝑗=1

where ℜ𝑗 is the rate of the j-th reaction. The kinetic scheme [45] used in the model was

developed for methane CPO on Rh and includes seven reactions: total oxidation of

methane; steam reforming of methane; water gas shift and reverse water gas shift;

96 KINETIC STUDY

methanation; H2 oxidation; CO oxidation. Since none of the original expressions for the

reaction rates has been used in this work, the seven-reaction kinetic scheme will not be

reported here for the sake of brevity. It is only worth underlining that WGS and rWGS

are expressed through two distinct expressions, since it was proven that it is not possible

to use one single equation to describe both the direct and the reverse steps.

In the case of CO oxidation in the absence of hydrogen, one single reaction is to be

considered: CO oxidation. In the case of CO PrOx, four of the original seven reactions

are in principle to be taken into account: the two oxidations, the water gas shift reaction

and the reverse water gas shift.

All the reaction rates had to be derived basically from scratch: the details of the

derivation are reported in the following paragraphs.

5.3 STUDY OF CO OXIDATION IN THE ABSENCE OF HYDROGEN

5.3.1 Introduction

Before performing a kinetic study on the oxidation of carbon monoxide in a hydrogen-

rich environment, an attempt was made to derive a kinetic expression for the oxidation

of CO alone. The effect of the concentration of CO and oxygen on the kinetics was

investigated through a number of tests performed at a GHSV=500000 NL/h/kg, in the

100-300 °C temperature range. The experiments have been discussed in depth in 4.3.

First, a differential analysis was performed by using low-conversion data to obtain a

power-law like expression for the reaction rate:

ℜ = 𝑘𝑃𝐶𝑂𝛼 𝑃𝑂2

𝛽

Then, starting from the results of the differential analysis, a more elaborate expression

was developed by exploiting the mathematical model of the annular reactor: the kinetic

parameters were properly modified by comparing the results of the model to the

experimental data.

STUDY OF CO OXIDATION IN THE ABSENCE OF HYDROGEN 97

5.3.2 Differential analysis

The data obtained for the effect of CO concentration, gathered by temperature and at a

constant partial pressure of oxygen, are reported in Table 2.1. ℜ is expressed in mol/s/kg,

while 𝑃𝐶𝑂 is expressed in atm.

Temperature Tests 𝒚𝑪𝑶 𝒚𝑶𝟐 𝒍𝒏(𝕽) 𝒍𝒏(𝑷𝑪𝑶)

180 °C Prova 43

Prova 47

Prova 48

0.01

0.02

0.005

0.01

0.01

0.01

-19.74

-21.38

-18.53

-4.61

-3.91

-5.30

200 °C Prova 43

Prova 47

0.01

0.005

0.01

0.01

-15.68

-16.57

-4.61

-3.91

Table 5.1: Data at varying CO concentration for the differential analysis in the absence of hydrogen.

By plotting the data in a bilogarithmic plot, Figure 5.1 is obtained:

Figure 5.1: Bilogarithmic plot for the data at varying CO concentration for the differential analysis.

ln(ℜ) = ln(𝑘𝐶𝑂) + 𝛼ln(𝑃𝐶𝑂)

Temperature Trendline R2 𝜶

180 °C y = -0.482x - 14.189 0.9925 -0.482

200 °C y = -0.7761x - 16.773 1 -0.776

Table 5.2: Results of the differential analysis on the reaction order with respect to CO.

-5,5

-5

-4,5

-4

-3,5

-3

-2,5

-2

- 1 9 - 1 8 - 1 7 - 1 6 - 1 5 - 1 4

LOG

(R)

LOG(P_CO)

EFFECT OF CO CONCENTRATION

180 °C 200 °C Lineare (180 °C) Lineare (200 °C)

98 KINETIC STUDY

Due to the very limited amount of data which could be considered suitable for a

differential analysis, the values derived for 𝛼 at 180 °C and 200 °C are quite different.

However, it is clear that the reaction rate depends on the concentration of CO with a

negative reaction order, possibly close to -0.6 (the average value is -0.629).

A similar analysis can be carried out for the effect of oxygen concentration. Here are

reported the experimental data:

Temperature Tests 𝒚𝑪𝑶 𝒚𝑶𝟐 𝒍𝒏(𝕽) 𝒍𝒏(𝑷𝑶𝟐)

180 °C Prova 43

Prova 44

Prova 45

Prova 46

0.01

0.01

0.01

0.01

0.01

0.0075

0.0025

0.02

-19.74

-20.06

-20.72

-18.22

-4.61

-4.89

-5.99

-3.91

200 °C Prova 43

Prova 44

Prova 45

0.01

0.01

0.01

0.01

0.0075

0.0025

-18.34

-18.76

-20.10

-4.61

-4.89

-5.99

220 °C Prova 40

Prova 44

Prova 45

0.01

0.01

0.01

0.005

0.0075

0.0025

-18.54

-18.12

-19.53

-5.30

-4.89

-5.99

Table 5.3: Data at varying O2 concentration for the differential analysis in the absence of hydrogen.

Figure 5.2: Bilogarithmic plot for the data at varying O2 concentration for the differential analysis.

-6,5

-6

-5,5

-5

-4,5

-4

-3,5

-3

- 1 8 , 5 - 1 8 - 1 7 , 5 - 1 7 - 1 6 , 5 - 1 6 - 1 5 , 5 - 1 5

LOG

(R)

LOG(P_O2)

EFFECT OF O2 CONCENTRATION

180 °C 200 °C 220 °C

Lineare (180 °C) Lineare (200 °C) Lineare (220 °C)


The results are the following:

ln(ℜ) = ln(𝑘𝑂2) + 𝛽ln(𝑃𝑂2)

Temperature Trendline R2 𝜷

180 °C y = 0.7635x + 10.18 0.869 0.764

200 °C y = 0.7951x + 9.9981 0.9991 0.795

220 °C y = 0.7683x + 8.9943 0.9929 0.768

Table 5.4: Results of the differential analysis on the reaction order with respect to O2.

In this case, the results are in very good agreement with each other. The reaction order

with respect to oxygen is positive, and its value is close to 0.8 (the average value is 0.776).

Thus, the analysis of single orders suggests a power-law expression for the oxidation of

CO in the absence of hydrogen as the following:

ℜ = 𝑘𝑃𝐶𝑂−0.63𝑃𝑂2

0.78

The result is in line with the literature.

5.3.3 Integration of CO oxidation into the model of the annular reactor

First of all, all the reactions but CO oxidation were removed from the original model of

the annular reactor, by zeroing their kinetic constant. Then, the kinetics of CO oxidation

was properly modified.

A power-law expression like the one obtained through the differential analysis would

have been inadequate for the simulation of the whole conversion range. In fact, since the

reaction rate is expressed as:

ℜ = 𝑘𝑃𝑂2𝛽

𝑃𝐶𝑂𝛼

as the conversion of CO tends to 1, the reaction rate tends to infinite, leading to numerical

issues. Thus, instead of using a simple expression like the one above, an expression based

on a reaction path was developed.

It is clear from the experimental results that, at least on a Pt-type catalyst, the reaction is

inhibited by CO and enhanced by oxygen: thus, a negative reaction order can be

expected for the first species, and a positive one for the second. This is confirmed in the

literature for the so-called low-rate branch. CO oxidation is indeed characterized by two

100 KINETIC STUDY

different reaction regimes: in the low-rate branch regime, occurring at low λ = 2𝑃𝑂2/𝑃𝐶𝑂

values and/or low temperatures, the surface is almost entirely covered by CO. In this

case, the reaction orders are close to -1 and +1 for CO and oxygen, respectively. On the

contrary, the high rate branch is characterized by a low CO surface coverage and occurs

at high temperatures, and/or high λ values: in this case, the reaction orders for CO and

oxygen are near +1 and 0, respectively [22].

Different Langmuir-Hinshelwood [46] expressions have been proposed in the literature,

with the surface reaction between CO and oxygen representing the rate-determining

step. If the adsorption of oxygen is assumed to be dissociative, the surface reaction

involves adsorbed carbon monoxide and atomic oxygen:

CO + * ↔ CO*

O2 + 2 * ↔ 2O*

CO* + O* → CO2 + *

leading to the following expression for the reaction rate:

ℜ =𝑘𝑃𝐶𝑂𝑃𝑂2

1/2

(1 + 𝐾𝐶𝑂𝑃𝐶𝑂 + √𝐾𝑂2𝑃𝑂2)2

On the contrary, if oxygen is assumed to adsorb without dissociating, the surface

reaction takes place between adsorbed carbon monoxide and molecular oxygen:

CO + * ↔ CO*

O2 + * ↔ O2*

CO* + O2* → CO2* + O*

leading to this other expression:


(1 + 𝐾𝐶𝑂𝑃𝐶𝑂 + 𝐾𝑂2𝑃𝑂2)2

The expression derived by assuming non-dissociative adsorption for oxygen was found

to better fit the experimental data. Hence, the following expression was implemented in

the reactor model:


(1 + 𝐾𝐶𝑂𝑃𝐶𝑂)2


where the term 𝐾𝑂2𝑃𝑂2 has been neglected. Indeed, the major inhibiting effect is the one

associated to carbon monoxide. Thus, the fraction of active sites occupied by species

other than CO can be assumed as negligible.

The parameter 𝑘 (which is actually a combination of constants) can be expressed as a

function of the temperature through modified Arrhenius’ law:

𝑘(𝑇) = 𝑘0exp [−𝐸𝑎𝑅(1

𝑇−

1

𝑇𝑟𝑒𝑓)]

In principle, 𝐾𝐶𝑂 should be considered a function of the temperature through a similar

expression:

𝐾𝐶𝑂(𝑇) = 𝐾𝐶𝑂,0exp [−∆𝐻𝑎𝑑𝑠,𝐶𝑂

𝑅(1

𝑇−

1

𝑇𝑟𝑒𝑓)]

where ∆𝐻𝑎𝑑𝑠,𝐶𝑂 itself has been proven to be a function of the CO coverage itself [47].

Thus, for the sake of simplicity, 𝐾𝐶𝑂 has been assumed constant.

By modelling different tests performed in the absence of hydrogen (see Figure 5.3, Figure

5.4 and Figure 5.5), the values for 𝑘0 and 𝐸𝑎 which better fit the data were estimated. The

kinetic parameters are reported in Table 5.5. A comparison between the model and the

reference test Prova 43 can be found in Figure 5.6.


(1 + 𝐾𝐶𝑂𝑃𝐶𝑂)2

𝒌𝟎

[mol/s/g_cat/atm1.5] 𝑬𝒂 [J/mol] 𝑻𝒓𝒆𝒇 [K] 𝑲𝒂𝒅𝒔,𝑪𝑶 [atm-1]

2.5E-00 73163 473 5.00E+02

Table 5.5: Kinetic parameters for CO oxidation in the absence of hydrogen.

102 KINETIC STUDY

100 150 200 250 300

0

20

40

60

80

100

GHSV = 300000 [NL/h/kg]

GHSV = 500000 [NL/h/kg]

GHSV = 1000000 [NL/h/kg]

CO

co

nve

rsio

n (

%)

Temperature (°C)

100 150 200 250 300

0

20

40

60

80

100

GHSV = 300000 [NL/h/kg]

GHSV = 500000 [NL/h/kg]

GHSV = 1000000 [NL/h/kg]

O2

co

nve

rsio

n (

%)

Temperature (°C)

Figure 5.3: Results of the model for the tests in the absence of hydrogen: effect of the GHSV.

100 150 200 250 300

0

20

40

60

80

100

CO 2%

CO 1%

CO 0.5%

CO

co

nvers

ion

(%

)

Temperature (°C)

50 100 150 200 250 300

0

20

40

60

80

100

CO 2%

CO 1%

CO 0.5%

O2 c

on

ve

rsio

n (

%)

Temperature (°C)

Figure 5.4: Results of the model for the tests in the absence of hydrogen: effect of CO concentration.

100 150 200 250 300

0

20

40

60

80

100

O2 1%

O2 0.75%

O2 0.25%

CO

co

nve

rsio

n (

%)

Temperature (°C)

100 150 200 250 300

0

20

40

60

80

100 O2 1%

O2 0.75%

O2 0.25%

O2

co

nve

rsio

n (

%)

Temperature (°C)

Figure 5.5: Results of the model for the tests in the absence of hydrogen: effect of O2 concentration.


100 150 200 250 300

0

20

40

60

80

100

Model

Experimental data (Prova 43)

CO

co

nve

rsio

n (

%)

Temperature (°C)

100 150 200 250 300

0

20

40

60

80

100 Model

Experimental data (Prova 43)

O2 c

onvers

ion (

%)

Temperature (°C)

Figure 5.6: Comparison between the model and Prova 43 (GHSV=500000 NL/h/kg, 1% CO, 1% O2, no hydrogen).

104 KINETIC STUDY

5.4 STUDY OF CO OXIDATION IN THE PRESENCE OF HYDROGEN

5.4.1 Preliminary considerations

In principle, a simple power-law expression for the oxidation of CO can be derived

starting from a differential analysis, without considering any other reaction taking place

in the system, as proposed by some authors [9], [23]. However, due to the limited amount

of data at low conversion of oxygen (which, in this case, is always to be considered the

limiting reactant due to the large excess of fuel), no valuable pieces of information could

be derived from a differential analysis. Still, it is possible to make some qualitative

considerations by observing the experimental data gathered at low temperatures which,

even if not rigorously exploitable for a differential analysis (especially because of the fact

that low-temperature data are the ones mostly affected by stabilization dynamics) are

representative of the trend of the reaction rates with the concentration of the reactants.

0 1 2 3 4

0,0014

0,0016

0,0018

0,0020

0,0022

0,0024

100 °C

120 °C

140 °C

yC

O2

% CO

0 1 2 3 40,0010

0,0015

0,0020

0,0025

0,0030

0,0035

0,0040 100 °C

120 °C

140 °C

yH

2O

% CO

Figure 5.7: Trends of yCO2 and yH2O as a function of the percentage of CO, at three different temperatures.

STUDY OF CO OXIDATION IN THE PRESENCE OF HYDROGEN 105

For instance, Figure 5.7 shows the trend of the concentration of CO2 and H2O as a

function of the percentage of CO at three different (nominal) temperatures. For

concentrations from 0.5% up to 2% CO, the trend of CO2 is decreasing with the amount

of CO. This is revealing a negative reaction order, possibly very close to -1, for the

oxidation of carbon monoxide with respect to CO, though data obtained at a 4% CO

concentration strongly deviate from this trend.

By observing the trend of yH2O as a function of CO concentration, it can be clearly seen

that the amount of water is decreasing as the concentration of CO increases, indicating

that this species also inhibits the kinetics of hydrogen oxidation. Again, this is true only

for concentrations from 0.5% up to 2%. Moreover, the amount of water produced at 0.5%

CO concentration is significantly higher, possibly indicating that the amount of CO is

too small to saturate the surface, as previously assumed (4.2.3).

The same analysis can be carried out for the effect of oxygen concentration (Figure 5.8).

In this case, the amount of both species increases more or less linearly with yO2. The

increase seems to be more than linear for a 2% and 4% oxygen concentration: however,

such high values of the stoichiometric ratio might be associated to a change in the

reaction mechanism: particularly low selectivities were also observed for such high

oxygen concentrations (Figure 4.8). Moreover, tests carried out at high oxygen

concentrations were characterized by significant deviations from the nominal

temperatures, and thus from the proper isothermal behaviour of the system.

To conclude, a negative reaction order with respect to CO, possibly close to -1, and a

positive reaction order with respect to oxygen, possibly close to +1, can be expected. CO

has been proven to be the must abundant species on the surface, almost saturating it at

least up to a certain temperature. Under the hypothesis of a competition between

hydrogen and CO for active sites, a negative reaction order with respect to CO for both

the kinetics suggests that the presence of a CO adlayer might be hindering oxygen

adsorption, which might thus represent the rate-determining step of the process.

106 KINETIC STUDY

0 1 2 3 40,0000

0,0005

0,0010

0,0015

0,0020

0,0025

0,0030

100 °C

120 °C

140 °C

yC

O2

% O2

0 1 2 3 40,000

0,001

0,002

0,003

0,004

0,005

0,006

0,007

0,008

0,009

100 °C

120 °C

140 °C

yH

2O

% O2

Figure 5.8: Trends of yCO2 and yH2O as a function of the percentage of O2, at three different temperatures.

Other considerations can be made by observing the Arrhenius’ plots in the presence and

in the absence of hydrogen. In order to build an Arrhenius plot, a pseudo-first-order

kinetics with respect to CO can be assumed:

ℜ𝐶𝑂 = −𝑘𝑎𝑝𝑝𝑃𝐶𝑂

𝑘𝑎𝑝𝑝 can be derived starting from the equation for a PFR:

𝑑�̇�𝐶𝑂𝑑𝑊𝑐𝑎𝑡

= ℜ𝐶𝑂

By assuming a constant volumetric flow rate,

𝑄𝑑𝑐𝐶𝑂𝑑𝑊𝑐𝑎𝑡

= ℜ𝐶𝑂


𝑄𝑑𝑐𝐶𝑂𝑑𝑊𝑐𝑎𝑡

= −𝑘𝑎𝑝𝑝𝑐𝐶𝑂𝑅𝑇

By integrating on the total mass of catalyst, the following expression is obtained:

through which 𝑘𝑎𝑝𝑝 can be calculated at a given temperature. Since:

𝑘𝑎𝑝𝑝 = 𝑘0,𝑎𝑝𝑝exp(−𝐸𝑎,𝑎𝑝𝑝

𝑅𝑇)

In logarithmic terms, the following equation for a straight line is obtained:

ln(𝑘𝑎𝑝𝑝) = ln(𝑘0,𝑎𝑝𝑝) −𝐸𝑎,𝑎𝑝𝑝

𝑅

1

𝑇

After interpolating the data through a linear regression, the apparent activation energy

of the reaction can be estimated from the slope of the straight line on the bilogarithmic

plot. The graphs were built using the data gathered for the tests at varying GHSV (4.2.2

and 4.3.2), and by referring to the nominal temperature.

Significant differences are present between the Arrhenius’ plots for CO oxidation in the

absence and in the presence of hydrogen. The Arrhenius’ plot for CO oxidation in

nitrogen (Figure 5.9) includes only data gathered above 200 °C (the reaction basically

does not start at lower temperatures). By plotting ln(𝑘𝑎𝑝𝑝) as a function of 1/T, no

change in the slope can be observed. In particular, the slope of the line intercepting the

data is associated to an average apparent activation energy of 110 kJ/mol.

Figure 5.9: Arrhenius' plot for CO oxidation in the absence of hydrogen.

y = -13231x + 23,925R² = 0,9759

-4,5

-4

-3,5

-3

-2,5

-2

-1,5

-1

-0,5

0

0,5

0 , 0 0 1 7 0 , 0 0 1 8 0 , 0 0 1 9 0 , 0 0 2 0 , 0 0 2 1 0 , 0 0 2 2

LN(K

_AP

P)

1/T [1/K]

NO HYDROGEN

108 KINETIC STUDY

The Arrhenius’ plot for CO oxidation in the presence of hydrogen (Figure 5.10) is

significantly different. First of all, it includes data gathered even at 120 °C. Moreover,

the graph for ln(𝑘𝑎𝑝𝑝) is characterized by the presence of two regions, with two different

slopes, thus indicating a possible change in the reaction mechanism.

Figure 5.10: Arrhenius' plot for CO oxidation in the presence of hydrogen.

Figure 5.11: Arrhenius' plots for CO oxidation in the presence of hydrogen. Left: low temperatures. Right: high

temperatures.

The apparent activation energies associated to the two low and high temperature range

are 20.9 kJ/mol and 49.8 kJ/mol, respectively.

-6

-5

-4

-3

-2

-1

0

0 , 0 0 1 5 0 , 0 0 1 7 0 , 0 0 1 9 0 , 0 0 2 1 0 , 0 0 2 3 0 , 0 0 2 5 0 , 0 0 2 7 0 , 0 0 2 9

LN(K

_AP

P)

1/T [1/K]

WITH HYDROGEN

y = -2508,4x + 2,3879R² = 0,9394

-4,3

-4,1

-3,9

-3,7

-3,5

-3,3

-3,1

-2,9

0 , 0 0 2 1 0 , 0 0 2 2 0 , 0 0 2 3 0 , 0 0 2 4 0 , 0 0 2 5 0 , 0 0 2 6

LN(K

_AP

P)

1/T [1/K]

y = -5985,4x + 9,9938R² = 0,9723

-4

-3,5

-3

-2,5

-2

-1,5

-1

-0,5

0 , 0 0 1 8 0 , 0 0 1 9 0 , 0 0 2 0 , 0 0 2 1 0 , 0 0 2 2 0 , 0 0 2 3

LN(K

_AP

P)

1/T [1/K]


Surprisingly enough, the activation energy is lower at lower temperatures. Thus, its

variation is probably associated to a change in the reaction mechanism and not to the

rise of diffusional limitations. Moreover, it can be clearly observed that the presence of

hydrogen is associated to a significant decrease in the activation energy of the process,

both at high and low temperatures, indicating that the presence of hydrogen might lead

to a change in the nature of the transition state of CO oxidation.

5.4.2 Choice of a reaction scheme

Differently from the case of CO oxidation alone, a reaction scheme for the oxidation of

CO in the presence of hydrogen requires at least two reactions.

As previously stated, the presence of hydrogen is indeed on the one hand detrimental

for the selectivity, but it strongly enhances the reactivity of CO at the same time,

suggesting a possible hydrogen-aided mechanism for the oxidation of carbon monoxide.

The following overall stoichiometry may be proposed, representing a concertated

mechanism for the oxidation of both hydrogen and carbon monoxide:

CO + H2 + O2 → CO2 + H2O

However, this expression would always lead to a 50% selectivity. Even if the

experimental data confirm that the selectivity is indeed rather close to 50% under most

operating conditions, strong deviations from this value can be observed for large oxygen

and CO concentrations (as it can be seen in the graphs for the selectivity in Figure 4.7

and Figure 4.8). Thus, a co-oxidation alone is not enough to describe the reacting system:

at least one independent oxidation reaction should be taken into account.

An alternative reaction scheme is the following:

H2 + 0.5 O2 → H2O

CO + H2O ↔ CO2 + H2

This scheme has been developed starting from some considerations about the nature of

hydrogen and CO oxidation. Hydrogen oxidation on Pt-type catalysts is known to

proceed very fast up to total conversion, even at room temperature. On the contrary, as

it could be see in the experiments carried out in the absence of hydrogen, CO oxidation

starts only around 200 °C. In the presence of hydrogen, the oxidation of H2 is somehow

hindered, while the reactivity of CO is strongly enhanced. Thus, one might assume that

110 KINETIC STUDY

the oxidation of CO takes place only as a hydrogen-mediated, indirect oxidation: the

water gas shift reaction is representative for this general class of reactions.

However, this reaction scheme seems not to explain two experimental observations.

First, the overcoming of the thermodynamic equilibrium at the higher temperatures,

which must be related to the presence of a parallel fast direct CO oxidation reaction.

Second, the fact that the selectivity seems to be independent on the GHSV (Figure 4.4),

excluding the possibility of an in-series mechanism. Thus, the contribution of a water

gas shift-like reaction is not to exclude, but it is probably an additional contribution

backing the two parallel oxidations.

Hence, neither a co-oxidation, nor water gas shift alone are enough to explain the

reactivity of CO in the presence of hydrogen, at least at the higher temperatures. The

following reaction scheme was thus considered:

CO2 + 0.5 O2 → CO2

H2 + 0.5 O2 → H2O

CO2 + H2 → CO + H2O

assuming that the two direct oxidations are taking place in parallel, one independently

from the other. The kinetics of both reactions will be, in principle, different from the one

in a hydrogen-free, or in a CO-free system. The reverse water gas shift is necessary to

account for the presence of the equilibrium constraint at higher temperatures, leading to

the consumption of CO2 as oxygen is depleted.

It should be underlined that the water gas shift reaction might indeed play some role at

low temperatures inside the real system. However, since no tests were carried out to

investigate the activity of the PrOx catalyst towards water gas shift in the 100-300 °C

temperature range, its role remains uncertain.

5.4.3 Integration of CO oxidation into the model of the annular reactor

For CO oxidation, an expression similar to the one derived in the absence of hydrogen

was selected:

ℜ𝐶𝑂 =𝑘𝐶𝑂𝑃𝐶𝑂𝑃𝑂2(1 + 𝛼𝑃𝐶𝑂)

2


where 𝛼 is a term accounting for the CO-related inhibition. The parameters were

adjusted by comparing the model to the experimental data. More in particular, it was

observed that the reaction seems to be inhibited by CO not as strongly as in the absence

of hydrogen. In principle, 𝛼 could be assumed to be a function of the temperature:

however, the introduction of a temperature dependence into the model did not change

significantly the quality of the results.

With regard to the oxidation of hydrogen, a similar expression was selected:

ℜ𝐻2 =𝑘𝐻2𝑃𝑂21 + 𝛼𝑃𝐶𝑂

where the term related to CO inhibition is still present. In theory, the partial pressure of

hydrogen should appear in the expression, as well: however, since the amount of

hydrogen is more or less constant throughout the reaction, it can be neglected.

While the two-reaction scheme, with the proper set of parameters, seemed to acceptably

fit most of the experimental data, still no maximum was present in the curve for CO

conversion. This proved that reverse water gas shift is essential in properly describing

the reactivity of the system. A simple kinetic expression for reverse water gas shift was

thus implemented into the model:

ℜ𝑟𝑊𝐺𝑆 = 𝑘𝑟𝑊𝐺𝑆𝑃𝐶𝑂2(1 − 𝜂𝑟𝑊𝐺𝑆)

where the term

𝜂𝑟𝑊𝐺𝑆 =𝐾𝑝

𝐾𝑒𝑞=

∏ 𝑃𝑖𝜈𝑖𝑁𝐶

𝑖=1

exp (−∆𝐺𝑅

0

𝑅𝑇)

is required to account for the fact that the reaction is taking place only if

thermodynamically favoured, i.e. if 𝐾𝑝 > 𝐾𝑒𝑞. Otherwise, the direct reaction prevails. At

equilibrium, 𝜂𝑟𝑊𝐺𝑆=1 and the reaction rate is null.

The parameters which better fit the experimental data are reported in Table 5.6. When

performing the fitting, low-temperature data (below 180 °C) were neglected. Indeed,

they showed low sensitivity to both temperature and concentration, and are possibly the

most affected by stabilization phenomena (4.2.1).

112 KINETIC STUDY

CO oxidation H2 oxidation rWGS

𝒌𝑪𝑶𝑷𝑪𝑶𝑷𝑶𝟐(𝟏 + 𝜶𝑷𝑪𝑶)

𝟐

𝒌𝑯𝟐𝑷𝑶𝟐𝟏 + 𝜶𝑷𝑪𝑶

𝒌𝒓𝑾𝑮𝑺𝑷𝑪𝑶𝟐

𝒌𝟎 [𝒎𝒐𝒍

𝒔𝒈𝒂𝒕𝒎𝒏] 1.5E+00 5.0E-03 4.0E-04

𝑬𝒂 [𝑱

𝒎𝒐𝒍] 49884 38244 99768

𝜶[𝒂𝒕𝒎−𝟏] 150 150 -

𝑻𝒓𝒆𝒇[𝑲] 473 473 553

Table 5.6: Kinetic parameters for CO oxidation in the presence of hydrogen.

100 150 200 250 300

0

20

40

60

80

100





CO

co

nve

rsio

n (

%)

Temperature (°C) 100 150 200 250 300

0

20

40

60

80

100

O2

co

nve

rsio

n (

%)

Temperature (°C)

100 200 300

0,000

0,002

0,004

0,006

0,008

0,010

0,012

0,014

yH

2O

Temperature (°C)

Figure 5.12: Results of the model. Effect of the GHSV.


The results obtained at different GHSVs are shown in Figure 5.12. The model agrees

fairly well with the data, even if the trend of water concentration does not perfectly

reflect the experimental one, especially at GHSV=300000 NL/h/kg.

The results of the simulations at different CO concentration are shown in Figure 5.13:

data obtained at 0.5% CO concentration have been excluded from the representation,

due to poor agreement with the experimental data. As already explained, these data

showed a much different trend with respect to the 1-4% concentration range, possibly

due to a non-saturated surface.

Finally, the results obtained at different O2 concentration can be found in Figure 5.14.

Data at 2% and 4% oxygen concentration, which as already discussed show a distinctly

different trend, have also been excluded: the amount of water was especially

underestimated.

100 150 200 250 300

0

20

40

60

80

100 CO 4%

CO 2%

CO 1%

CO

co

nve

rsio

n (

%)

Temperature (°C) 100 150 200 250 300

0

20

40

60

80

100

O2

co

nve

rsio

n (

%)

Temperature (°C)

100 200 300

0,000

0,002

0,004

0,006

0,008

0,010

0,012

0,014

yH

2O

Temperature (°C)

Figure 5.13: Results of the model. Effect of CO concentration.

114 KINETIC STUDY

100 150 200 250 300

0

20

40

60

80

100 O2 1%

O2 0.75%

O2 0.5%

O2 0.25%

CO

con

vers

ion

(%

)

Temperature (°C) 100 150 200 250 300

0

20

40

60

80

100

O2

co

nve

rsio

n (

%)

Temperature (°C)

100 150 200 250 300

0,000

0,002

0,004

0,006

0,008

0,010

0,012

0,014

yH

2O

Temperature (°C)

Figure 5.14: Results of the model. Effect of O2 concentration.


5.4.4 Comparison with methanation

The preferential oxidation of carbon monoxide presents some drawbacks. In order to

minimize the conversion of hydrogen, the amount of oxygen fed to a CO preferential

oxidation unit should be controlled very carefully. The mixing phase is also problematic

in terms of safety. Moreover, the reactor should be able of operating in a wide

temperature range to guarantee safe operation [48].

A selective methanation (CO-SMET) step could in principle replace the preferential

oxidation: no additional reactant is required, and the process is also inherently easier to

control, thanks to the lower exothermicity of methanation with respect to the oxidations

of CO and hydrogen. Hence, a comparison between CO PrOx and selective methanation

was performed by comparing the reaction rate of PrOx to the one of methanation under

the same operating conditions.

The rate of the methanation reaction can be expressed with the following expression,

derived within a previous Thesis work [30]:

𝑟𝑀𝐸𝑇 =𝛼𝑀𝐸𝑇𝑃𝐶𝑂𝑃𝐻2(1 + 𝐾𝐶𝑂𝑃𝐶𝑂)

2

where

𝛼𝑀𝐸𝑇(𝑇) = 𝛼0𝑒𝑥𝑝 [−𝐸𝑎𝑅(1

𝑇−

1

𝑇𝑟𝑒𝑓)]

𝐾𝐶𝑂(𝑇) = 𝐾𝐶𝑂,0exp [−∆𝐻𝑎𝑑𝑠,𝐶𝑂

𝑅(1

𝑇−

1

𝑇𝑟𝑒𝑓)]

𝜶𝟎

[mol/s/kg/atm2] 𝑬𝒂 [kJ/mol] 𝑲𝑪𝑶,𝟎 [atm-1]

∆𝑯𝒂𝒅𝒔,𝑪𝑶

[kJ/mol] 𝑻𝒓𝒆𝒇 [K]

1.69 23.81 493.35 -46.95 493.15

Table 5.7: Kinetic parameters for methanation (from [30]).

116 KINETIC STUDY

The rate of the two reactions can be estimated for a reference composition of 40% H2, 1%

CO and, in the case of CO PrOx, 1% O2, at temperatures ranging from 200 to 300 °C.

T [°C] ℜMET [mol/s/kg] ℜCO OX [mol/s/kg]

200 0.00007 0.02400

220 0.00019 0.04014

240 0.00049 0.06449

260 0.00109 0.10000

280 0.00216 0.15021

300 0.00381 0.21933

Table 5.8: Comparison between the initial rates of methanation and PrOx.

As it can be observed in Table 5.8, the rate of CO oxidation is for each temperature at

least two orders of magnitude higher than the one of methanation. This is of course only

an estimation of the initial rate, since the rate of the reaction decreases along the axial

coordinate, as the reactants are consumed. Moreover, the oxidation of CO is taking place

in parallel to hydrogen oxidation: modelling the whole reactor would provide more

meaningful results.

Still, the fact that the rate of CO oxidation is this larger with respect to the one of

methanation is significant. Indeed, a much smaller amount of catalyst should be

necessary in the case of the PrOx reactor to guarantee the desired outlet conversion,

making it a more compact solution, preferable to methanation.

117

CONCLUSIONS

The thorough removal of carbon monoxide is a critical issue in hydrogen-rich streams

fed to PEM fuel cells. Indeed, the presence of even small amounts of this species is

associated to efficiency loss and irreversible damage of the Pt anode due to CO-induced

poisoning. The reaction of preferential oxidation of CO (CO PrOx), long-established at

the industrial scale, is vastly employed for this purpose.

A kinetic study of the preferential oxidation reaction was carried out within this Thesis

work, both through experiments at the laboratory scale, and through a modelling phase.

The experiments were initially performed in a diluted packed bed reactor, by using the

catalyst in the form of a powder. Due to issues related to the exothermicity of the

reaction, the dilution ratio was increased more times. Still, radial temperature gradients

were likely affecting the results, according to Mears’ criteria on interparticle heat

transport limitations. This led to the choice of an alternative reactor configuration.

An annular reactor was used to carry out an in-depth kinetic study of CO PrOx. This

type of reactor is particularly indicated for fast, exothermic reactions, thanks to the

additional paths for heat dispersion and the possibility of operating at high space

velocities in the absence of pressure drops. Moreover, it allows to monitor the axial

temperature profile, and to check whether the reaction is actually taking place under

quasi-isothermal conditions.

The effect of different process parameters (space velocity, concentration of CO,

concentration of oxygen) was investigated in the annular system. Tests in the absence of

hydrogen were also performed, in order to evaluate its effect on the kinetics of CO

oxidation. As it was observed, hydrogen has on the one hand a negative impact on the

selectivity, but on the other hand it strongly enhances the reactivity of CO at low

temperatures. This is also proven by the decrease in the apparent activation energy of

the reaction in the presence of hydrogen.

During the experiments, the catalyst showed a very slow, but perfectly reversible

deactivation. In particular, the conversion of both CO and oxygen decreased with the

time on stream. This phenomenon is possibly related to the gradual covering of the

118 CONCLUSIONS

surface by carbon monoxide, producing a poisoning effect on the reactions. As CO

occupies more and more active sites, the adsorption of oxygen is hindered and so is the

oxidation of both CO and hydrogen.

The experimental results clearly showed that both the oxidations are inhibited by CO,

and are on the contrary favoured by large oxygen concentrations. Starting from these

qualitative considerations, simple rate expressions were developed for CO oxidation

(both in the presence and in the absence of hydrogen) and hydrogen oxidation. The

reverse water gas shift reaction, which takes place at high temperatures as oxygen is

depleted, was also added to the kinetic scheme. The parameters for the rate expressions

were obtained by comparing the experimental data to the results of a previously

developed 1-d heterogeneous model for the annular reactor.

The obtained rate expressions can be used for the simulation of a PrOx reactor. A

preliminary comparison with the estimated rate for CO methanation over a Rh catalyst

suggests that the removal of the residual amount of carbon monoxide through

preferential oxidation can be obtained by means of a smaller amount of catalyst and

larger GHSVs than in the case of hydrogenation to CH4.

120

BIBLIOGRAPHY

[1] A. Beretta, G. Groppi, C. Ribani, G. Fares and C. Tregambe, "Development of a

Catalytic Fuel Processor for a 10 kW Combined Heat and Power System:

Experimental and Modeling Analysis of the Steam Reforming Unit,"

ChemEngineering, 2018.

[2] F. Ullmann, Ullmann's Encyclopedia of Industrial Chemistry, Wiley-VCH.

[3] J. D. Holladay, J. Hu, D. L. King and Y. Wang, An overview of hydrogen production

technologies, 2009.

[4] A. F. Ghenciu, "Review of fuel processing catalysts for hydrogen production in

PEM fuel cell systems," Current Opinion in Solid State and Materials Science, 2002.

[5] R. Farrauto, Y. Liu, W. Ruettinger, O. Ilinich, L. Shore and T. Giroux, "Precious

metal catalysts supported on ceramic and metal monolithic structures for the

hydrogen economy," Catalysis Reviews - Science and Engineering, 2007.

[6] D. B. Pal, R. Chand, S. N. Upadhyay and P. K. Mishra, Performance of water gas shift

reaction catalysts: A review, 2018.

[7] L. M. Martínez T, O. H. Laguna, C. López-Cartes and M. A. Centeno, "Synthesis

and characterization of Rh/MnO2-CeO2/Al2O3 catalysts for CO-PrOx reaction,"

Molecular Catalysis, 2017.

[8] M. L. Brown, Jr. and A. W. Green, "Treatment Of Gases". Patent 3 088 919, 1963.

[9] F. Romero-Sarria, S. Garcia-Dali, S. Palma, E. M. Jimenez-Barrera, L. Oliviero, P.

Bazin and J. A. Odriozola, "The role of carbon overlayers on Pt-based catalysts for

H2-cleanup by CO-PROX," Surface Science, 2016.

[10] P. C. Hulteberg, J. G. Brandin, F. A. Silversand and M. Lundberg, "Preferential

oxidation of carbon monoxide on mounted and unmounted noble-metal catalysts

in hydrogen-rich streams," International Journal of Hydrogen Energy, 2005.

121

[11] Y. Choi and H. G. Stenger, "Kinetics, simulation and insights for CO selective

oxidation in fuel cell applications," Journal of Power Sources, 2004.

[12] K. Liu, A. Wang and T. Zhang, Recent advances in preferential oxidation of co reaction

over platinum group metal catalysts, 2012.

[13] Y. F. Han, M. J. Kahlich, M. Kinne and R. J. Behm, "CO removal from realistic

methanol reformate via preferential oxidation - Performance of a Rh/MgO catalyst

and comparison to Ru/γ-Al 2O3, and Pt/γ-Al2O3," Applied Catalysis B:

Environmental, 2004.

[14] F. Mariño, C. Descorme and D. Duprez, "Noble metal catalysts for the preferential

oxidation of carbon monoxide in the presence of hydrogen (PROX)," Applied

Catalysis B: Environmental, 2004.

[15] O. Korotkikh and R. Farrauto, "Selective catalytic oxidation of CO in H2: Fuel cell

applications," Catalysis Today, 2000.

[16] A. Sirijaruphan, J. G. Goodwin and R. W. Rice, "Effect of Fe promotion on the

surface reaction parameters of Pt/γ-Al2O3 for the selective oxidation of CO,"

Journal of Catalysis, 2004.

[17] X. Liu, O. Korotkikh and R. Farrauto, "Selective catalytic oxidation of CO in H2:

Structural study of Fe oxide-promoted Pt/alumina catalyst," Applied Catalysis A:

General, 2002.

[18] C. Pedrero, T. Waku and E. Iglesia, "Oxidation of CO in H2-CO mixtures catalyzed

by platinum: Alkali effects on rates and selectivity," Journal of Catalysis, 2005.

[19] M. S. Moreno, E. López, M. E. Adrover, N. J. Divins and J. Llorca, "CO-PrOx over

nano-Au/TiO2: Monolithic catalyst performance and empirical kinetic model

fitting," International Journal of Hydrogen Energy, 2016.

[20] H. C. Lee and D. H. Kim, "Kinetics of CO and H2 oxidation over CuO-CeO2 catalyst

in H2 mixtures with CO2 and H2O," Catalysis Today, 2008.

122

[21] G. Avgouropoulos, T. Ioannides, C. Papadopoulou, J. Batista, S. Hocevar and H. K.

Matralis, "A comparative study of Pt/γ-Al2O3, Au/α-Fe2O3 and CuO-CeO2

catalysts for the selective oxidation of carbon monoxide in excess hydrogen," in

Catalysis Today, 2002.

[22] M. J. Kahlich, H. A. Gasteiger and R. J. Behm, "Kinetics of the selective CO oxidation

in H2-rich gas on Pt/Al2O3," Journal of Catalysis, 1997.

[23] D. H. Kim and M. S. Lim, "Kinetics of selective CO oxidation in hydrogen-rich

mixtures on Pt/alumina catalysts," Applied Catalysis A: General, 2002.

[24] E. J. Bissett, S. H. Oh and R. M. Sinkevitch, "Pt surface kinetics for a PrOx reactor

for fuel cell feedstream processing," Chemical Engineering Science, vol. 60, no. 17, pp.

4709-4721, 9 2005.

[25] W. Hauptmann, M. Votsmeier, H. Vogel and D. G. Vlachos, "Modeling the

simultaneous oxidation of CO and H2 on Pt - Promoting effect of H2 on the CO-

light-off," Applied Catalysis A: General, 2011.

[26] S. Salomons, R. E. Hayes and M. Votsmeier, "The promotion of carbon monoxide

oxidation by hydrogen on supported platinum catalyst," Applied Catalysis A:

General, 2009.

[27] S. Salomons, M. Votsmeier, R. E. Hayes, A. Drochner, H. Vogel and J. Gieshof, "CO

and H2 oxidation on a platinum monolith diesel oxidation catalyst," Catalysis Today,

2006.

[28] M. M. Schubert, M. J. Kahlich, H. A. Gasteiger and R. J. Behm, "Correlation between

CO surface coverage and selectivity/kinetics for the preferential CO oxidation over

Pt/γ-Al2O3 and Au/α-Fe2O3: an in-situ DRIFTS study," 1999.

[29] A. B. Mhadeshwar and D. G. Vlachos, "Hierarchical, multiscale surface reaction

mechanism development: CO and H2 oxidation, water-gas shift, and preferential

oxidation of CO on Rh," Journal of Catalysis, 2005.

[30] R. Rinaldi and V. Troisi, "Small scale fuel processing: a kinetic investigation of CO

methanation," Politecnico di Milano, 2018.

123

[31] C. Cristiani, M. Valentini, M. Merazzi, S. Neglia and P. Forzatti, "Effect of ageing

time on chemical and rheological evolution in γ-Al2O3 slurries for dip-coating," in

Catalysis Today, 2005.

[32] A. G. Bayram and V. Piazza, "Syngas production from ethanol and formic acid over

Rh/Al2O3: a kinetic study in annular reactor," Politecnico di Milano, 2019.

[33] Colorado State University, "Chemical Equilibrium Calculator," [Online]. Available:

http://navier.engr.colostate.edu/code/code-4/index.html.

[34] R. Rota, Fondamenti di termodinamica dell'ingegneria chimica, Pitagora Editrice,

2015.

[35] J. Pérez-Ramírez, R. J. Berger, G. Mul, F. Kapteijn and J. A. Moulijn, "Six-flow

reactor technology a review on fast catalyst screening and kinetic studies," Catalysis

Today, 2000.

[36] A. Beretta, G. Groppi, L. Majocchi and P. Forzatti, "Potentialities and draw-backs

of the experimental approach to the study of high T and high GHSV kinetics,"

Applied Catalysis A: General, 1999.

[37] C. Beniani, "Experimental and modelling analysis of the water gas shift reactor in

a combined heat and power unit," Politecnico di Milano, 2018.

[38] R. H. Nibbelke, M. A. Campman, J. H. Hoebink and G. B. Marin, "Kinetic study of

the CO oxidation over Pt/γ-Al2O3 and Pt/Rh/CeO2/γ-Al2O3 in the presence of

H2O and CO2," Journal of Catalysis, 1997.

[39] A. Sirijaruphan, J. G. Goodwin and R. W. Rice, "Investigation of the initial rapid

deactivation of platinum catalysts during the selective oxidation of carbon

monoxide," Journal of Catalysis, vol. 221, no. 2, pp. 288-293, 25 1 2004.

[40] D. E. Mears, "Tests for Transport Limitations in Experimental Catalytic Reactors,"

Industrial and Engineering Chemistry Process Design and Development, 1971.

[41] "EUROKIN spreadsheet on requirements for measurement of intrinsic kinetics in

the gas-solid fixed-bed reactor," 2012. [Online]. Available:

124

https://www.eurokin.org/wp-content/uploads/webtool/EUROKIN_fixed-

bed_html.htm.

[42] "ASALI - CatalyticFOAM," [Online]. Available:

http://www.catalyticfoam.polimi.it/asali-2/.

[43] A. Beretta, P. Baiardi, D. Prina and P. Forzatti, "Analysis of a catalytic annular

reactor for very short contact times," Chemical Engineering Science, 1999.

[44] G. F. Froment, K. B. Bischoff and J. De Wilde, Chemical Reactor Analysis and

Design - 3rd Ed., Wiley, 2011.

[45] A. Donazzi, A. Beretta, G. Groppi and P. Forzatti, "Catalytic partial oxidation of

methane over a 4% Rh/α-Al2O3 catalyst. Part I: Kinetic study in annular reactor,"

Journal of Catalysis, 2008.

[46] B. W. Wojciechowski and S. P. Asprey, "Kinetic studies using temperature-

scanning: The oxidation of carbon monoxide," Applied Catalysis A: General, 2000.

[47] A. Bourane, O. Dulaurent and D. Bianchi, "Heats of adsorption of linear and

multibound adsorbed CO species on a Pt/Al2O3 catalyst using in situ infrared

spectroscopy under adsorption equilibrium," Journal of Catalysis, 2000.

[48] M. A. Ashraf, G. Ercolino, S. Specchia and V. Specchia, "Final step for CO syngas

clean-up: Comparison between CO-PROX and CO-SMET processes," in

International Journal of Hydrogen Energy, 2014.

small-scale fuel processing: kinetic study of co

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