Kinetic Modeling of the Catalytic Steam Reforming of High-Density

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Kinetic modelling of the catalytic steam reforming of HDPE pyrolysis volatiles Itsaso Barbarias, Gartzen Lopez, Maite Artetxe, Aitor Arregi, Javier Bilbao, and Martin Olazar Energy Fuels, Just Accepted Manuscript • DOI: 10.1021/acs.energyfuels.7b01909 • Publication Date (Web): 02 Oct 2017 Downloaded from http://pubs.acs.org on October 3, 2017

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Kinetic modelling of the catalytic steam reforming of HDPE

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pyrolysis volatiles

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Itsaso Barbarias, Gartzen Lopez*, Maite Artetxe, Aitor, Arregi, Javier Bilbao and

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Martin Olazar

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Department of Chemical Engineering University of the Basque Country UPV/EHU,

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P.O. Box 644-E48080 Bilbao (Spain). [email protected]

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Abstract

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The kinetics of the steam reforming of HDPE pyrolysis volatiles is studied on a Ni

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commercial catalyst in a fluidized bed reactor in-line with the pyrolysis reactor (a

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conical spouted bed reactor at 500 ºC). Steam reforming reactions have been carried out

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under the following conditions: 600-700 ºC and space time 0-16.7 gcat min gHDPE-1.

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Based on the composition of HDPE pyrolysis volatiles, a kinetic scheme is assumed

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with four reactions (C5+ hydrocarbon reforming, C2-C4 hydrocarbon reforming, CH4

14

reforming and WGS reaction). Moreover, the kinetics of the deactivation has been

15

quantified with an expression dependant on C5+ hydrocarbon concentration (coke

16

precursors). Calculation of the kinetic parameters is conducted by nonlinear multiple

17

regression, fitting the experimental data to those calculated by the mass conservation

18

equations for each component in the reaction medium. The overall kinetic model

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proposed describes accurately the evolution of the main products (H2, CO2, CO) with

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time on stream in the range of operating conditions studied.

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Keywords: kinetic model; steam reforming; plastics; pyrolysis; conical spouted bed; Ni

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catalyst; hydrogen

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1.

Introduction

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The challenges involving the management of plastic wastes and their reduction should

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be directed towards their recycling and valorisation. However, the rate of increase in

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plastic wastes is higher than the rate of recycling, and plastic wastes are increasingly

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sent to landfill or incinerated 1. Fast pyrolysis of plastics is one of the most promising

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routes to obtain valuables products, such as gas or liquid fuels, monomers and

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chemicals 2-6. Furthermore, the pyrolysis process is versatile, environmentally friendly

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(absence of oxygen hinders the formation of hazardous volatile compounds) and allows

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treating different plastics (HDPE, PS, PET, PMMA, PP, PVC) 4, 7-18. Consequently,

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plastic pyrolysis has been widely studied and has reached a considerable level of

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development at both laboratory scale and pilot plant based on several alternative

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technologies 3, 19, 20. However, there are many technological and economical factors that

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hinder the scalability of the process, such as the great diversity of plastic types, their

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additive content, energy requirements, collection and classification costs, and product

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valorisation and commercialisation 4.

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From a technological front of view, the greatest difficulty of plastic pyrolysis process is

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the sticky nature of fused plastic particles, which causes particle agglomeration and the

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subsequent defluidization in fluidized bed reactors and poor heat transfer rates in other

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technologies. A previous work has proven that the conical spouted bed reactor (CSBR)

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has suitable features for minimizing these problems, and therefore for operating with

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continuous feed and at a low temperature (500 ºC) without temperature gradients in the

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reactor and without bed defluidization problems 9, 11, 21. The plastic particles fed into the

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reactor rapidly melt and uniformly coat the sand particles in the bed due to their cyclic

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and vigorous movement, which in turn leads to high heat transfer rates between phases,

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and therefore facilitates a rapid volatilization of the plastics. Particle collisions with 2 ACS Paragon Plus Environment

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high momentum and the capacity of the spout to break the agglomerates contribute to

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avoiding the agglomeration of adhesive sand particles coated with fused plastic.

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Furthermore, the low residence time of the volatiles maximizes their yield with a

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uniform composition, which is essential for their in-line valorisation in a reforming

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reactor.

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Some studies focus on combining thermal process (pyrolysis, gasification) with

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catalytic steam reforming to obtain H2 from plastics waste 22-27. Wu and Williams 28

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developed a pyrolysis-gasification process from polypropylene using a two fixed bed

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reaction system and they investigated the influence of different catalysts on H2

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concentration. Czernik and French 29 proved the efficiency of the pyrolysis-steam

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reforming process to obtain H2 from polyethylene and polypropylene using two

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fluidized bed reactors. The use of two steps has the following advantages 30: i)

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temperature in each step can be optimized separately; ii) the feed into the catalytic

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reactor is uniform and its composition established by the pyrolysis conditions; ii) the

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catalyst is more effective in the transformation of volatiles because particles are not

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coated with fused polymer and so all the catalyst is available for contacting with

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pyrolysis volatiles, thereby decreasing the amount of catalyst needed for obtaining the

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maximum yield of H2.

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Nickel based catalysts are the most used in a two-stage reaction system for obtaining H2

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from plastics 22, 24, 25, 31-34. Noble metals such as Ru and Rh are the most active ones 23,

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35

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French 29 achieved 80 % of the H2 yield allowable by stoichiometry in the pyrolysis-

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reforming of PP using a naphtha reforming commercial Ni catalyst. Barbarias et al. 31

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obtained higher H2 yields (92.5 %) in the pyrolysis-reforming of HDPE on a

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commercial CH4 reforming Ni catalyst, with the pyrolysis step being at 500 ºC and the

, but their high cost makes them inappropriate for waste valorisation. Czernik and


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reforming one at 700 ºC. Wu and Williams 36, using a two-stage fixed bed reactor,

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reported high H2 content in the gas obtained using Ni/CeO2 (75.5 vol.%) and Ni-Al (64

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vol.%). Furthermore, according to Li et al. 37 SiO2 is a more adequate support for Ni

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than TiO2, MgO, ZrO2 and Al2O3. However, Choudhary et al. 38 highlighted Ni/ZrO2

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and Ni/Ce-NaY catalysts based on their regeneration capability.

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Kinetic models for each step in the pyrolysis-catalytic steam reforming process are

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required for the simulation and optimization of operating conditions and subsequent

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scaling up of the process. Nevertheless, studies have only been conducted on plastic

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pyrolysis kinetics and most of them based on thermogravimetry 39-41. Thus, the lumped

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kinetic models proposed in the literature for the catalytic pyrolysis of waste plastics are

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scarce 42-44. Similarly, few detailed kinetic studies dealing with waste plastic

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gasification have been reported in the literature 45, 46. Thus, studies on the kinetic

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modelling of hydrocarbon steam reforming are almost exclusively limited to CH4

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reforming. The model by Numaguchi and Kikuchi 47 is based on reforming reactions

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and WGS, and the one by Xu and Froment 48 also consider the methanation reaction.

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In this paper, the kinetic modelling of the steam reforming of HDPE pyrolysis volatiles

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is addressed on a Ni based catalyst. The overall kinetic model has been developed by

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coupling the one corresponding to zero time on stream (fresh catalyst) and the one for

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catalyst deactivation. The global model proposed faithfully predicts product distribution

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in the reforming reactor and its evolution with time on stream.

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2.

Materials and methods

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2.1.

Catalyst

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A commercial Ni catalyst (G90LDP) provided by Süd Chemie and used in industry for

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CH4 reforming has been used. Its generic formulation is NiO (14%), CaAl3O4, Al2O3. 4 ACS Paragon Plus Environment

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The choice of this catalyst is motivated firstly in its availability without reproducibility

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problems, and secondly in its good performance in previous works (low deactivation by

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coke deposition and high capacity for hydrogen production) 49. The catalyst has been

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sieved to 0.4-0.8 mm, which is a suitable particle size to attain stable fluidization

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regime. Prior to the reforming reaction, the catalyst has been subjected to an in situ

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reduction process at 710 ºC for 4 h under a 10 vol.% H2 stream. The N2 adsorption-

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desorption isotherm of this catalyst has already reported elsewhere 30, 50, and shows low

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porosity with a BET surface area of 19 m2 g-1 and an average pore diameter of 122 Å.

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2.2.

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Fig. 1 shows the experimental bench scale plant used in the HDPE continuous

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pyrolysis-reforming runs. The HDPE (supplied by Dow Chemical) is fed continuously

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(0.75 g min-1) into the pyrolysis reactor by means of a solid feeding system. The

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pyrolysis reactor is a CSBR and its dimensions are established according to previous

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hydrodynamics studies 51 and based on the prior application of this reactor in the

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pyrolysis and gasification of biomass 52-54, plastics 8, 9, 55, 56, and tyres 57. The reactor is

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placed inside a radiant oven of 1250 W that heats the pyrolysis reactor and its lower

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section, which acts as a gas preheater and is filled with an inert ceramic material that

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increases the surface area for heat transfer. The bed is made up of 50 g of sand (0.3-

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0.35mm) and N2 is used to fluidize it during the heating process until the reaction

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temperature of 500 ºC.

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The water required in the reforming step, which is vaporized by means of an electrical

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cartridge, acts as fluidizing agent in both reactors (pyrolyser and reformer). Previous

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results confirmed that steam behaves as an inert compound in the pyrolysis of HDPE at

Experimental equipment and conditions


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moderate temperatures (500 ºC) 50, with the results being similar to those reported by

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Elordi et al. 9 using N2 as fluidizing agent.

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Figure 1

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The volatile stream obtained in the pyrolysis reactor at 500 ºC (constant temperature in

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all experiments) is mainly composed of waxes (C21+, 67 wt%), diesel fraction (C12-C20

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25.6 wt%), and gasoline fraction (C5-C11, 5.9 wt%) , with the remaining compounds

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being light hydrocarbons (C2-C4, 1.47 wt%) and CH4 (0.03 wt%) 21, 58. These volatiles

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are fed in-line into a fluidized bed reforming reactor (Fig. 1). The fluidized bed reactor

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has advantages over the fixed bed reactor, due to the isothermicity of the bed and its

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scalability, in particular for a process with catalyst deactivation, as it may be operated

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by circulating the catalyst for its regeneration.

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The experimental data for the kinetic study have been obtained under the following

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operating conditions: 600, 650 and 700 ºC; space-time, 2.1, 4.1, 8.3 12.5 and 16.7 gcat

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min gHDPE-1; time on stream, up to 130 min. The temperature range selected for the

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fitting of the kinetic model is based on the experience acquired in previous studies

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dealing with the reforming of plastic pyrolysis volatiles 31. Thus, operation at

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temperatures below 600 ºC is hindered by the low reforming kinetic rate and the low

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conversion of pyrolysis products. Similarly, operation above 700 ºC led to irreversible

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catalyst deactivation due to Ni sintering 49. The space-time has been modified varying

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the catalyst to sand ratio, but maintaining constant the total mass in the bed (25 g).

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Water flow rate was 3 mL min-1 in the different experiments performed, which

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corresponds to a steam flow rate of 3.73 NL min-1. Moreover, the continuous polymer

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feed rate to the pyrolysis reactor was of 0.75 g min-1, accordingly, steam/plastic ratio


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was of 4 and steam/carbon (S/C) ratio in the reforming step of 7 given that HDPE was

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completely converted into volatiles in the pyrolysis step.

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The volatile products circulate through a volatile condensation system and a sample is

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regularly analyzed by an on-line Varian 3900 chromatograph. The non-condensable

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gases are analyzed by micro GC Varian 4900.

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3.

Results

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3.1.

Kinetic model

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Figure 2 shows the kinetic scheme proposed considering 4 individual reactions,

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corresponding to the reforming reaction of C5+, hydrocarbons C2-C4 hydrocarbons, CH4

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and WGS reaction. The consideration of hydrocarbons in three different lumps (C5+

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hydrocarbons, C2-C4 hydrocarbons and CH4) is based on their different origin, kinetic

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behaviour and reactivity. Thus, CH4 is more stable and less reactive than the other

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hydrocarbons involved in the reforming of HDPE pyrolysis volatiles. C2-C4

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hydrocarbons showed a clearly more reactive behaviour than CH4, with their origin

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being the thermal cracking of C5+ hydrocarbons. Finally, C5+ lump is mainly made up of

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pyrolysis waxes and, to a lesser extent, the diesel fraction, with its reactivity being

159

similar to that of C2-C4 fraction.

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Coke formation reaction has not been considered in the reforming step modelling, given

161

that although its effect on deactivation is important, its content is small 31 and does not

162

have a significant impact on the overall mass balance of the reforming step.

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Figure 2

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The reactions corresponding to the proposed kinetic scheme are as follows, with eqs. 1

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and 2 written with generic stoichiometries: 7 ACS Paragon Plus Environment

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C5+ fraction reforming:

mC 5+ + nH 2 O → nH 2 + nCO

(1)

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C2-C4 fraction reforming:

m(C 2 − C 4 ) + nH 2 O → nH 2 + nCO

(2)

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CH4 reforming:

CH 4 + H 2O → 3H 2 + CO

(3)

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Water Gas Shift reaction (WGS):

CO + H 2O ↔ H 2 + CO2

(4)

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Ideal plug flow has been considered for gas circulation in the fluidized bed, i.e.,

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unidirectional flow without radial gradients of concentration. Moreover, isothermal bed

172

has been assumed, as temperature differences at different radial and longitudinal

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positions are lower than 1 ºC. The total molar flow rate, FT, varies along the reactor due

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to the increase in the number of moles in the reforming reactions. Therefore, the mass

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conservation equation in a differential element of catalyst mass (dW) for component i in

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the kinetic scheme in Fig. 2 is as follows:

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dFi d (FT ⋅ X i ) dX i dF = = FT + X i T = (ri )0 dW dW dW dW

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where W is the catalyst mass and Xi the molar fraction of component i, expressed on a

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wet basis.

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The formation rate of each component i at zero time on stream has been established

181

considering all the reaction steps in which it is involved:

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(ri )0 = ∑ (υi )j (rj )0

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where (υi)j is the stoichiometric coefficient of component i in the reaction j of the

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kinetic scheme, and (rj)0 is the rate of reaction j at zero time on stream. Given that

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hydrocarbon reactants C5+ and C2-C4 are compound mixtures, their average composition

(5)

j

(6)


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has been considered in order to determine the stoichiometric coefficient of the

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reforming reaction.

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The reaction rate expressions, (rj)0 in Eqs. (7)-(10), are formulated assuming they are

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first-order for each reactant. This assumption is based on the following considerations:

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i) a simpler model is attained by avoiding the incorporation of reaction orders as

191

adjustable parameters, ii) reaction kinetics has a clear physical meaning depending on

192

reactants partial pressure and iii) previous methane steam reforming kinetic studies

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showed that the reaction orders with respect to methane are around 1, even though those

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with respect to steam are more diverse 59-63. It is to note that lower reaction orders were

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reported in the steam reforming of heavier hydrocarbons such as toluene and benzene,

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being these results associated with their higher reactivity in relation to that of methane

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64, 65

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(r1)0 = k1 XC5+ XH2O

(7)

199

(r2)0 = k2 XC2-C4 XH2O

(8)

200

(r3)0 = k3 XCH4 XH2O

(9)

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(r4)0 = kWGS (XCO XH2O - XH2XCO2/KWGS)

202

where Xi is the molar fraction of each component in the reaction medium and kj is the

203

kinetic constant of the reaction step j in the kinetic scheme (Eqs. (1)-(4)).

204

The equilibrium constant of the WGS reaction has been calculated by the following

205

expression:

206

1 1   K = exp a + b + c log( T ) + dT + eT 2 + f 2  T T  

. The expressions are:

(10)

(11)


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where a-f parameters (Table 1) have been calculated by means of the methodology

208

described by Smith et al. 66.

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Table 1

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As aforementioned, the experimental results required for the modelling have been

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obtained by conducting pyrolysis-reforming of HDPE at three reforming temperatures

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(600, 650 and 700 ºC) and using five different catalyst masses at each temperature (1.5,

213

3.1, 6.3, 9.4 and 12.5 gcat), corresponding to the space-times of 2.1, 4.1, 8.3, 12.5 and

214

16.7 gcat min gHDPE-1, respectively. The concentration values obtained for the kinetic

215

model have been the molar fractions of H2, CO2, CO, H2O, and non-reacted CH4, C2-C4

216

and C5+ hydrocarbons. It should be noted that the runs have also been carried out

217

without catalyst (space time of 0 gcat min gHDPE-1), in which the high conversion by

218

thermal cracking has been proven. This thermal cracking is a fast process, and so occurs

219

close to the inlet of the catalytic reactor. Therefore, the composition of the reaction

220

medium at the inlet of the reforming reactor has been assumed to be the composition

221

resulting from the thermal cracking, and not that of the pyrolysis reactor outlet.

222

Furthermore, it has been confirmed that this consideration is required to attain a suitable

223

fit of the kinetic model.

224

3.2.

225

Given the lack of studies concerning the kinetic modelling of hydrocarbon reforming,

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the methodology developed by Oar-Arteta et al. 67 for the reforming of oxygenated

227

compounds (DME) has been considered as a reference. The general aspects of this study

228

have been developed based on the guidelines established for the kinetic modelling of

229

catalytic process at zero time on stream (without taking into account the deactivation)

230

updated by Toch et al. 68.

Methodology for the kinetic study


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The parameters of the kinetic model proposed have been calculated by nonlinear

232

multiple regression. The calculation consists in minimizing the objective function

233

defined as the weighted sum of squared differences between the experimental and

234

calculated values of molar fractions:

235

OF = ∑ w i φ i = ∑ w i ∑ (X ∗i , j − X i, j ) 2

nc

nc

p

i =1

i =1

j=1

(12)

236

in which wi is the weight factor for each component in the reaction scheme, φi is the

237

sum of squared residuals for each compound; X ∗i , j is the experimental value of the molar

238

fraction of component i under the experimental condition j; X i , j is the corresponding

239

value of molar fraction calculated solving the mass balance for component i, Eq. (5); nc

240

is the number of products in the reaction scheme; and p is the total number of

241

experimental conditions.

242

The parameters to be optimized are the kinetics constants of each reaction j, Eqs. (7)-

243

(10). To minimize the correlation between the pre-exponential factor and the activation

244

energy in the Arrhenius equation, the reparameterized form of this equation has been

245

adopted, Eq. (13), where k ∗j is the kinetic constant at a reference temperature, T* (650

246

ºC):

247

 E j  1 1  k j = k ∗j exp −  − ∗   R  T T 

248

Given that there are no available data of repeated experiments for an estimation of the

249

initial variances, the weight factors (Eq. (14)) are established as inversely proportional

250

to the average composition of each product in the range of operating conditions studied

251

69

(13)

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wi =

1

(14)

p

∑ Xi j=1

253

In order to calculate the molar fractions of the products, a calculation program has been

254

developed in Matlab. The main program, which acts as an input and output interface for

255

data and results, calls the subroutine for nonlinear regression (“fminsearch” function in

256

Matlab). This subroutine minimizes the error between the experimental data and those

257

predicted by the kinetic model proposed. Matlab ode15s function has been used to

258

integrate the differential equations.

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3.3.

260

Table 2 shows the kinetic parameters (kinetic constants at the reference temperature and

261

activation energies) of the rate equations described above, Eqs. (7)-(10), obtained by

262

fitting the experimental results to the overall kinetic model. In addition, Table 2 also

263

shows the corresponding sums of squared residuals. As observed, the kinetic constant

264

for C5+ hydrocarbons is higher than those for the remaining ones, i.e., it is twice that for

265

the reforming reaction of C2-C4 hydrocarbons (k*1 > k*2). Reforming of CH4, however,

266

is less favoured and slower than the reforming of C5+ and C2-C4 hydrocarbons.

267

Accordingly, the reforming kinetic constants follow then subsequent order: k*1 > k*2 >

268

k*3. This result reveals that the reforming reaction rate is favoured as the carbon content

269

in the hydrocarbon is higher, and therefore CH4 is the most refractory to reform

270

followed by C2-C4 hydrocarbons and C5+ hydrocarbons. The aforementioned results

271

justify the consideration of three lumps to describe the reactions undergone by the

272

hydrocarbons formed in the reforming of HDPE pyrolysis volatiles. The kinetic

273

constant of the WGS reaction (kWGS = 3.0·10-1 mol gcat-1 min-1) reveals that this reaction

274

is the fastest one.

Kinetic model fitting and kinetic parameters


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Table 2

276

The adequacy of the fit is shown in Fig. 3, in which the experimental evolution of

277

products molar fractions (points) with catalyst mass (or space time) are compared with

278

those calculated with the kinetic model proposed (lines), at 600 ºC (Fig. 3a,b), 650 ºC

279

(Fig. 3c,d) and 700 ºC (Fig. 3e,f). As aforementioned, the concentration values for 0 gcat

280

min gHDPE-1 space time are those corresponding to the thermal cracking at the reaction

281

temperature considered (experiments without catalyst). As observed in Fig. 3, the fit

282

between experimental and calculated data is adequate and the model proposed predicts

283

reasonably well the effect temperature and catalyst mass (or space-time) have on

284

product distribution.

285

Steam reforming reactions, Eqs. (1)-(3), are endothermic, and therefore an increase in

286

temperature favours the reforming reactions of C5+ and C2-C4 hydrocarbons and CH4,

287

whose concentrations decrease with space time more sharply as temperature is higher

288

(Figures 3b, 3d and 3f). A favourable effect of temperature on H2 concentration is

289

observed when comparing these concentrations at 600 ºC (Fig. 3a) and 650 ºC (Fig. 3b).

290

A further increase to 700 ºC does not increase H2 concentration because the WGS

291

reaction (reversible and exothermic) is less favoured.

292

Figure 3

293

Furthermore, at 600 ºC (Fig. 3a) H2 concentration increases when catalyst mass (and

294

therefore space-time) is increased to a value of 9.4 g (12 gcat min gHDPE-1) due to the

295

enhancement of reforming reactions and the WGS reaction. Accordingly, full reforming

296

of C5+ hydrocarbons is obtained with a space-time of 12 gcat min gHDPE-1. At this

297

temperature, the amount of catalyst has a significant effect on product distribution; that


Energy & Fuels

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298

is, increasing the amount of catalyst the yields of H2, CO and CO2 increase, whereas the

299

one of H2O decreases.

300

At 650 and 700 ºC (Fig. 3d and 3f, respectively) the yields of hydrocarbon fractions also

301

decrease as space time is increased, and full conversion of C5+ hydrocarbons is obtained

302

above 6 gcat min gHDPE-1. At 700 ºC, C2-C4 hydrocarbons are almost fully converted, but

303

CH4 hardly reaches full conversion. Furthermore, as observed in Fig. 3c and 3e,

304

corresponding to 650 ºC and 700 ºC, equilibrium conversion is reached in the WGS

305

reaction for 6 gcat min gHDPE-1, and H2, CO2 and CO concentrations keep constant for

306

higher space times.

307

These trends with temperature and space-time on product distribution are similar to

308

those obtained by other authors using Ni based catalyst in the reforming step in

309

fluidized bed reactors 24, 29 or fixed beds reactors 22, 23, 32, 50.

310

3.4.

311

Given the relevance of catalyst deactivation by coke in the hydrocarbon reforming

312

process on Ni catalysts, it is mandatory to include the deactivation equation in the

313

kinetic model in order to quantify the evolution of product concentrations with time on

314

stream. Previous studied showed that deactivation is mainly caused by the blockage of

315

Ni active sites by polyaromatics structures, with the main precursors for their formation

316

being the wax chains in the C5+ fraction 49. Therefore, the deactivation has been

317

quantified based on a kinetic equation established considering deactivation rate

318

proportional to C5+ hydrocarbon concentration in the reaction medium:

319

−

Kinetic modelling of the deactivation

da = k d X C5+ a dt

(15)


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320

where XC5+ is the molar fraction of C5+ hydrocarbons and kd the kinetic constant of

321

deactivation. This activity has been defined as the ratio between the reaction rate at a

322

specific time on stream and the reaction rate at zero time on stream:

323

a=

324

It should be noted that non-selective deactivation has been assumed, and hence the same

325

activity has been considered for the different steps of the kinetic scheme in Fig. 2

326

(reforming of C5+ hydrocarbons, C2-C4 hydrocarbons and CH4, and WGS reaction).

327

According to Eq. (16), the formation rate of component i at time on stream t is

328

calculated by multiplying the formation rate at zero time on stream (ri)0, Eqs. (7)-(10),

329

by activity. Thus, considering deactivation, the reaction rates in the model are as

330

follows:

331

r1 = k1 XC5+ XH2O a

(17)

332

r2 = k2 XC2-C4 XH2O a

(18)

333

r3 = k3 XCH4 XH2O a

(19)

334

r4 = kWGS (XCO XH2O - XH2 XCO2/KWGS)·a

(20)

335

The equilibrium constant of the WGS reaction has been calculated by means of Eq. (11)

336

using the parameters in Table 1.

337

The calculation of the kinetic constant of deactivation (which has been also

338

reparameterized) follows a similar procedure to that described for the kinetic parameters

339

at zero time on stream. Thus, the error function, Eq. (12), has been minimized by fitting

340

the experimental values of molar fractions of the components in the reaction medium to

341

the values corresponding to the same time on stream calculated by integrating Eqs. (17)-

ri

(16)

(ri )0


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342

(20) together with the deactivation equation, Eq. (15). In this calculation, the kinetic

343

constants previously determined for zero time on stream (Table 2) have been used.

344

The value calculated for the kinetic constant of deactivation at the reference temperature

345

(650 ºC) and its corresponding activation energy have been 2.2 min-1 and 10.8 kJ mol-1,

346

respectively. Fig. 4 shows the evolution of catalyst activity with time on stream,

347

calculated with Eq. (15) for the three temperatures studied, and a space time of 16.7 gcat

348

min gHDPE-1 taken as an example. It should be pointed out that, although the deactivation

349

kinetic constant increases with temperature, the decrease in catalyst activity is lower as

350

the reforming temperature is increased. This result is evidence of the importance of

351

considering the C5+ hydrocarbon concentration (coke precursor concentration in the

352

reaction medium) in the kinetic equation of deactivation, Eq. (15).

353

Figure 4

354

The fitting of the overall kinetic model (considering deactivation kinetics) is shown in

355

Fig. 5, in which the experimental data (points) are compared with model predictions

356

(lines) at the three temperatures, 600 ºC (Fig. 5a), 650 ºC (Fig. 5b) and 700 ºC (Fig. 5c).

357

As observed, the proposed model suitably predicts the evolutions with time on stream of

358

the molar fractions of H2, CO2, CO and H2O.

359

Figure 5

360

It should be noted that catalyst deactivation attenuates as temperature is raised and, at a

361

reforming temperature of 700 ºC (Figure 5c), the catalyst is only slightly deactivated

362

after 120 min on stream. The explanation lies in the fact that an increase in temperature

363

enhances the reforming reactions kinetics, which is consistent with the model proposed.

364

Thus, the minor presence of C5+ hydrocarbons (with a high content of waxes) in the

365

reaction medium, which are the coke precursors, is responsible for the decrease in coke 16 ACS Paragon Plus Environment

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366

deposition rate, and therefore for the lower catalyst deactivation rate. Consequently, it

367

may be concluded that 700 ºC is the most suitable reaction temperature.

368

4.

369

The kinetic model proposed describes the evolution with time on stream of products

370

distribution (H2, CO, CO2, H2O) in the steam reforming of HDPE pyrolysis volatiles, in

371

a wide range of operating conditions (temperature from 600 to 700 ºC, space time from

372

0 to 16.7 gcat min gHDPE-1). In spite of the simplicity of the kinetic approach used, the

373

model developed satisfactorily describes the reforming of HDPE derived hydrocarbons

374

both at zero time on stream and under deactivation conditions.

375

In order to quantify the effect of catalyst deactivation by coke, a kinetic equation of

376

deactivation has been established, which accounts for the deactivation caused by C5+

377

hydrocarbons at the reactor inlet. The results (described by the model) show that the

378

temperature of 700 ºC is the most suitable for obtaining the maximum yield and

379

concentration of H2, with the values being 85.7 wt% and 70 vol%, respectively, for a

380

space time of 16.7 gcat min gHDPE-1. Under these conditions the deactivation of the

381

catalyst is moderate.

382

The overall kinetic model proposed is an essential tool for simulating reaction-

383

regeneration strategies, aimed to the scaling up of pyrolysis-reforming systems for the

384

upgrading of plastics wastes.

385

5.

Nomenclature

a

activity

Ej, Ej*

activation energy of step j in the kinetic scheme and its corresponding

Conclusions


Energy & Fuels

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18 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

value at the reference temperature (kJ mol-1)

FO

error objective function

Fi, FT

molar flow rate of compound i and total molar flow rate (mol min-1)

kd

kinetic constant of deactivation (min-1)

kj, kj*

kinetic constant of step j in the kinetic scheme and its corresponding value at the reference temperature (mol gcat-1 min-1)

KWGS

equilibrium constant of the WGS reaction

nc

number of compounds in the kinetic scheme

p

total number of experimental conditions

R

constant of gases (kJ mol-1 K-1)

(ri)0, (rj)0

reaction rate of each compound i formation and of the step j at zero time on stream (mol gcat-1 min-1)

ri, rj

reaction rate of each compound i formation and of the step j (mol gcat-1 min-1)

SSE

sum of squared errors

t

time (min)

T, T*

temperature and reference temperature (K)

W

catalyst mass (gcat)

wi

weight factor for each compound i

Xi

molar fraction of compound i in the reaction medium (on a wet basis)

Xi,j*, Xi,j

experimental and calculated molar fractions (on a wet basis) of


Page 19 of 43

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compound i in the step j in the reaction medium

(υi)j

stoichiometric coefficient of compound i in the step j

φi

sum of squared residuals for each compound i

386

Acknowledgments

387

This work was carried out with financial support from the Ministry of Economy and

388

Competitiveness of the Spanish Government (CTQ2016-75535-R (AEI/FEDER, UE)

389

and CTQ2014-59574-JIN (AEI/FEDER, UE)), the Basque Government (IT748-13) and

390

the University of the Basque Country (UFI 11/39). I. Barbarias thanks the University of

391

the Basque Country for her postgraduate grant (UPV/EHU 2016).

392

References

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59. Ahmed, K.; Foger, K. Kinetics of internal steam reforming of methane on Ni/YSZ-

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based anodes for solid oxide fuel cells. Catal. Today 2000, 63, 479-487.

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60. Brus, G. Experimental and numerical studies on chemically reacting gas flow in the

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porous structure of a solid oxide fuel cells internal fuel reformer. Int. J. Hydrogen

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Energy 2012, 37, 17225-17234.

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61. Anna Sciazko, Yosuke Komatsu, Grzegorz Brus, Shinji Kimijima and Janusz S.

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Szmyd. A novel approach to the experimental study on methane/steam reforming

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kinetics using the Orthogonal Least Squares method. J. Power Sources 2014, 262, 245-

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254.

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62. King, D. L.; Strohm, J. J.; Wang, X. Q.; Roh, H. -.; Wang, C.; Chin, Y. -.; Wang,

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Y.; Lin, Y.; Rozmiarek, R.; Singh, P. Effect of nickel microstructure on methane steam-

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reforming activity of Ni-YSZ cermet anode catalyst. J. Catal. 2008, 258, 356-365.

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63. Timmermann, H.; Fouquet, D.; Weber, A.; Ivers-Tiffée, E.; Hennings, U.; Reimert,

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R. Internal reforming of methane at Ni/YSZ and Ni/CGO SOFC cermet anodes. Fuel

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Cells 2006, 6, 307-313.


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64. Koike, M.; Li, D.; Watanabe, H.; Nakagawa, Y.; Tomishige, K. Comparative study

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on steam reforming of model aromatic compounds of biomass tar over Ni and Ni-Fe

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alloy nanoparticles. Appl Catal A 2015, 506, 151-162.

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65. Tomishige, K.; Li, D.; Tamura, M.; Nakagawa, Y. Nickel-iron alloy catalysts for

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reforming of hydrocarbons: preparation, structure, and catalytic properties. Catal. Sci.

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Technol. 2017, 7, 3952-3979.

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66. Smith, J. M. Introduction to Chemical Engineering Thermodynamics, 7th Edition.

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McGraw-Hill 2005.

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67. Oar-Arteta, L.; Aguayo, A. T.; Remiro, A.; Arandia, A.; Bilbao, J.; Gayubo, A. G.

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Kinetics of the steam reforming of dimethyl ether over CuFe2O4/γ-Al2O3. Chem. Eng.

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J. 2016, 306, 401-412.

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68. Toch, K.; Thybaut, J. W.; Marin, G. B. A systematic methodology for kinetic

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modeling of chemical reactions applied to n-hexane hydroisomerization. AIChE J. 2015,

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69. Gayubo, A. G.; Alonso, A.; Valle, B.; Aguayo, A. T.; Olazar, M.; Bilbao, J. Kinetic

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modelling for the transformation of bioethanol into olefins on a hydrothermally stable

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Ni–HZSM-5 catalyst considering the deactivation by coke. Chem. Eng. J. 2011, 167,

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262-277.

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589

Figure captions

590

Figure 1.

Scheme of the laboratory scale catalytic steam reforming plant.

591

Figure 2.

Kinetic scheme proposed for the reforming of the volatiles formed in the pyrolysis of HDPE.

592

593

Figure 3.

Comparison of experimental (points) and calculated (lines) values for the

594

products molar fractions at zero time on stream, for different catalyst

595

masses, at 600 ºC (a,b), 650 ºC (c,d) and 700 ºC (e,f).

596

Figure 4.

temperatures studied.

597

598

Evolution of catalyst activity with time on stream at the three

Figure 5.

Comparison of the experimental results for the evolution of products

599

molar fractions obtained using different catalyst masses (points) with

600

those predicted using the model (lines) at 600 ºC (a), 650 ºC (b) and 700

601

ºC (c).

602


Energy & Fuels

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603

Table 1. Parameters of the equilibrium constant in the WGS reaction.

KWGS

a

b

c

d

e

f

-1.8·101

5.8·103

1.8

-2.7·10-4

0.0

-5.8·104

604 605


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606

Table 2. Kinetics parameters of best fit for the kinetic model proposed. Parameters k1*, mol gcat-1 min-1

2.0·10-1

k2*, mol gcat-1 min-1

1.2·10-1

k3*, mol gcat-1 min-1

6.7·10-2

kWGS*, mol gcat-1 min-1

3.0·10-1

E1, kJ mol-1

17.0

E2, kJ mol-1

17.1

E3, kJ mol-1

30.4

EWGS, kJ mol-1

32.5

SSE

0.086

607

608


Energy & Fuels

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Figure 1. Scheme of the laboratory scale catalytic steam reforming plant. 548x459mm (96 x 96 DPI)

ACS Paragon Plus Environment

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Figure 2. Kinetic scheme proposed for the reforming of the volatiles formed in the pyrolysis of HDPE. 413x331mm (96 x 96 DPI)


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1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Figure 3. Comparison of experimental (points) and calculated (lines) values for the products molar fractions at zero time on stream, for different catalyst masses, at 600 ºC (a,b), 650 ºC (c,d) and 700 ºC (e,f). 104x104mm (600 x 600 DPI)


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1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Figure 4. Evolution of catalyst activity with time on stream at the three temperatures studied. 104x104mm (600 x 600 DPI)


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Figure 5. Comparison of the experimental results for the evolution of products molar fractions obtained using different catalyst masses (points) with those predicted using the model (lines) at 600 ºC (a), 650 ºC (b) and 700 ºC (c). 104x104mm (600 x 600 DPI)


Energy & Fuels

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Kinetic Modeling of the Catalytic Steam Reforming of High-Density

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