CFD investigation of hydrodynamics, heat transfer and cracking reactions in a large-scale fluidized catalytic cracking riser

Accepted Manuscript

CFD investigation of hydrodynamics, heat transfer and cracking reactions in a large-scale fluidized catalytic cracking riser Qi Yang , Abdallah S. Berrouk , Yupeng Du , Hui Zhao , Chaohe Yang , Mohammad Abdur Rakib , Abdulhamid Mohamed , Anood Taher PII: DOI: Reference:

S0307-904X(16)30327-4 10.1016/j.apm.2016.06.016 APM 11221

To appear in:

Applied Mathematical Modelling

Received date: Revised date: Accepted date:

22 September 2015 23 May 2016 16 June 2016

Please cite this article as: Qi Yang , Abdallah S. Berrouk , Yupeng Du , Hui Zhao , Chaohe Yang , Mohammad Abdur Rakib , Abdulhamid Mohamed , Anood Taher , CFD investigation of hydrodynamics, heat transfer and cracking reactions in a large-scale fluidized catalytic cracking riser, Applied Mathematical Modelling (2016), doi: 10.1016/j.apm.2016.06.016

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Highlights Hydrodynamics coupled with cracking reactions in an FCC riser were studied.

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Two-fluid mathematical approach was used to model the fluid-particle flow.

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Intricate phenomena near a series of injecting nozzles were revealed in detail.

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A recommendation of an appropriate increasing of the feed injection angle was proposed.

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Yields and reaction rates of main products and catalyst activity were investigated.

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CFD investigation of hydrodynamics, heat transfer

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and cracking reactions in a large-scale fluidized catalytic cracking riser

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Qi Yang1, 2, Abdallah S. Berrouk1,*, Yupeng Du2, Hui Zhao2, Chaohe Yang2, Mohammad Abdur Rakib3, Abdulhamid Mohamed3, Anood Taher3 1

State Key Laboratory of Heavy Oil Processing, China University of Petroleum, Qingdao 266580, China

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Chemical Engineering Department, Petroleum Institute, Abu Dhabi 2533, United Arab Emirates

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Takreer Research Centre, Abu Dhabi 3593, United Arab Emirates

*Corresponding author:

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Dr Abdallah S. Berrouk Chemical Engineering Department Petroleum Institute, Abu Dhabi, UAE Phone: +971 26075408 Email: [email protected]

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ABSTRACT A three-dimensional reactive gas-particle CFD model was built to study the hydrodynamics, heat transfer and cracking reaction behaviors within an industrial Fluid Catalytic Cracking (FCC)

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riser reactor designed to maximize propylene production. The two-fluid methodology (TFM) was used to simulate the riser hydrodynamics with solid phase properties derived from the kinetic theory of granular flows (KTGF). An 11-lump kinetic model was selected to represent the cracking reaction network in the CFD model. The selection of the kinetic model is dictated by

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the properties of the feedstock processed and the aim of the process which is maximizing propylene. A novel treatment of the coke component was conducted by incorporating coke into the secondary granular phase which is more realistic since carbon deposition occurs on catalyst

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phase. Momentum transfer, heat transfer and reaction behavior inside the riser were discussed in detail and inhomogeneity in these aspects were observed especially above the high speed

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injection nozzles. The numerical results of this investigation show a good agreement with the process real data on the yield distribution despite the use of a coarse grid to mesh such an

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industrial scale FCC riser. The methodology employed used and the results obtained should

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serve as guidelines for possible process redesign and optimization.

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Keywords: fluidized catalytic cracking, riser, flow-reaction model, numerical simulation, injection angle

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1. Introduction Fluidized catalytic cracking (FCC) is the most commonly-used converting process of heavy petroleum fractions worldwide. Commercially useful products such as gasoline, diesel, middle distillates and light olefins such as propylene can be obtained from less valuable heavy oil using

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the FCC process. Typically, an FCC unit consists of two main parts; the riser reactor and the regenerator. The former hosts the cracking reactions that yield the desirable products and the latter regenerate the catalyst particles used to crack the heavy oil in the riser. An FCC riser is

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generally divided into four parts; the prelifting zone, the feedstock injection zone, the full reaction zone and the quenching zone. [1] FCC risers have been commercialized for over 70 years with many innovative technologies being implemented to satisfy evolving market demands and meet new challenges linked to the processing of an increasingly heavier crude oil. Some of

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these emerging FCC technologies such as atmospheric residue maximizing gas and gasoline (ARGG) process, maximizing iso-paraffin in cracked naphtha (MIP) process, [2] riser reaction

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termination technology to minimize over cracking, [3] mixing of deactivated catalysts with

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regenerated catalysts [4] to achieve low temperature contact with catalysts, two stage riser fluidized catalytic cracking for maximizing propylene (TSRFCC-TMP), [5,6] and flexible dual-

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riser fluid catalytic cracking (FDFCC) [7] process to decrease gasoline olefin and increase propylene, have been widely adopted. Recently, a noticeable increase in the global demand for

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propylene has shifted the focus of refineries towards FCC technologies that maximize production of propylene in order to achieve economic profits. Several recent studies [1,8-10] have explored the design scheme, preparation of catalysts, and corresponding lumped kinetic models in FCC riser aimed at maximizing propylene. In the two stage riser catalytic pyrolysis for maximizing propylene yield (TMP) [6] technology, atmospheric residue and C4 mixture gas are processed in

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the first stage of the riser, whereas recycle oil and light gasoline are fed into the second stage. C4 mixture gas and light gasoline contribute a great deal to the promotion of propylene production. [11] Both C4 mixture gas and light gasoline prefer high reaction temperature and high catalystto-oil ratio whilst atmospheric and recycle oil require low temperature and short reaction time.

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To achieve this duality, a fuel-stratified injection technique is introduced to first bring C4 mixture gas/light gasoline into contact with high temperature regenerated catalyst and consequentially cool down hot catalyst. Then, atmospheric residue/recycle oil can go through low-temperature cracking reactions on the surface of the cooled catalysts and stop the cracking process of

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previous C4 mixture gas/light gasoline due to their stronger absorption ability. This technology is reported to have: (i) expanded residuum refining capacity, (ii) bettered recovery rate of light crude, and (iii) increased diesel and propylene yield. [5] Another FCC process named flexible

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dual-riser fluid catalytic cracking (FDFCC) containing two risers is designed to minimize olefin in cracked gasoline and improve propylene yield which can compensate the gasoline loss. [7,12]

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As heavy oil cracking reaction and naphtha upgrading require different operating conditions, a second riser is often added to provide an independent zone for naphtha upgrading. Optimal

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operating conditions and residence time are necessary in both risers in order to attenuate mutual

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interruption between heavy oil cracking and naphtha upgrading, besides, increase the final propylene yield.

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FCC technology has been evolving steadily over the last couple of decades thanks to the significant amount of research that has been dedicated to study the complex flow system taking place in riser reactors. Indeed, the interrelated and simultaneous hydrodynamics, heat transfer, and cracking reactions have made detailed experimental investigations in industrial risers unachievable. Robust and reliable experimental instruments have also been lacking to measure

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reactive multiphase systems characterized by complex hydrodynamics, high temperature and significant numbers of produced species. To circumvent this lack of detailed knowledge on this important process to the refining industry, researchers have turned to state-of-the-art computational techniques such as Computational Fluid Dynamics (CFD) in order to gain critical

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insight into the intricate phenomena occurring inside FCC riser reactors. As a result, a wide range of studies that involve numerical simulations of hydrodynamics, heat transfer and cracking reactions in FCC risers have been conducted. Theologos and Markatos [13] established a 3D CFD model that incorporated two-phase flow, interphase heat transfer, and a 3-lump kinetic

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model (a kinetic model based on three lumped components) to describe the cracking reactions. The impact of feed injector geometry on hydrodynamics and cracking process at the FCC riser bottom was explored. In another work, they used a 10-lump kinetic model to closely investigate

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the performance of the whole riser as function of the number and geometry of feed injectors. [14] It was reported that selectivity of primary products was improved by increasing the number of

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feed injection nozzles from three to twelve. Gao et al. [15] also conducted a 3D reactive twophase-flow numerical experiment to evaluate the performance of FCC risers using a 13-lump

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kinetic model. Simulation results suggested that reacting flow patterns inside the riser were

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extremely complex especially in the feed injection zone where high injecting velocities prevail. Chang et al. [14, 16] pointed out, in similar works using 12-lump kinetic model, that a high

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temperature, long reaction time combined with a low catalyst-to-oil ratio should benefit the production of low-olefin clean gasoline and propylene. In a different work, Li et al. [17] reported the effect of nozzle jet velocity, nozzle position and nozzle angle on the performance of feed injection zone at the bottom of the riser. Nozzle angle played a significant role in the feed mixing zone while nozzle position contributed a little. It was highlighted that nozzle angle larger than

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30° was preferable. Other recent works that used the CFD technique have reported interesting results on the flow system and chemistry in FCC risers [18-19]. Chen et al. demonstrated that an increase in feed spray velocity would facilitate the feed diffusion and consequently enhance gasoil conversion. However, the excessively increased jet velocity intensified back-mixing near

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the riser wall and catalyst breakage [20]. Xiong et al. [21] were the first to establish a catalyst property that was included in a 6-lump kinetic model. The latter considered three key catalyst parameters: zeolite to matrix specific surface area ratio (Z/M), total Bronsted to Lewis acid amount ratio (B/L) and zeolite unit cell size (UCS) in the simulation of catalytic cracking

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conversion inside FCC risers. The latter was used by Lopes et al. [22] to find out numerically that small changes in the riser outlet configuration impose a significant effect on the flow patterns and cracking product yields. Using CFD, Alvarez-Castro et al. [23] evaluated the

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performance of a FCC riser based on product yield distribution and explored optimal conditions favoring a better flow homogeneity and attenuating unwanted core-annulus flow structure that

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often hinders gasoil conversion. Li et al. [24] utilized the CFD methodology to show that the electrostatics had a weak influence upon the particle concentration, gas-phase temperature and

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product yield distribution in a reactive FCC riser.

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In the current work, hydrodynamics, heat transfer and cracking reactions in the first riser of a commercial large-scale two-riser FCC unit were investigated using a 3D two-phase reactive flow

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model based on the Eulerian-Eulerian two-fluid methodology (TFM). The CFD model incorporated an 11-lump kinetic model [25] to demonstrate different mechanisms of maximizing propylene. The choice of the 11-lump kinetic model was dictated by the properties of the feedstock processed by the simulated riser and the aim of the process which was maximizing propylene. The geometry of the commercial riser reactor was considered with all its complexity

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as opposed to previous numerical investigations which used simplified geometries in order to easily explore optimal operating conditions. The riser geometry considered herein includes three groups of nozzles and two diameter enlargement sections. The simulated riser is a large-scale industrial reactor with the processing capacity of 6×106 ton/year. This riser is characterized by

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large dimensional scale differences between the largest diameter of the riser (over 2500 mm) and that of the nozzle (less than 100 mm). The large scale geometry and large dimensional scale differences have made hard to achieve both grid and numerical convergences. Furthermore, accounting for the mass transfer between the two phases was another challenge to address. This

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was due to the fact that the reaction source term of coke was connected to the catalyst phase which is more realistic given that coke deposits on catalyst particles as portrayed by the Coke on Catalyst (COC) deactivation model. All of the above mentioned challenges were addressed in the

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best way possible to strike the right balance between a reasonable execution time and good accuracy for such a large scale industrial case where the interplay between hydrodynamics, heat

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transfer and reaction kinetics is very complicated. The latter was discussed in detail in this work and the simulated product distribution was compared with the process data of the investigated

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FCC riser and used as a criterion for grid convergence.

2. Process and hardware description

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The large scale FCC riser simulated in this work is part of a dual-riser industrial FCC unit.

Feedstock is injected into the lower part of both risers. The first riser, which is the subject of this study, receives atmospheric residue and aims at maximizing propylene while the second riser, which is an auxiliary riser, is fed with recycle oil (unconverted oil from the first riser) to improve the overall conversion of the FCC unit. Figure 1 shows the geometry of the first riser (the

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detailed dimensions of the riser are confidential to the technology supplier). Pre-lifting steam is fed to the bottom of the riser while regenerated catalysts flow from an oblique downward conveying pipe into the riser at a higher position. As pre-lifting steam and catalysts move upward along the FCC riser axial direction, they encounter the feedstock which is injected with

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atomizing steam through a series of feed oil injection nozzles. The latter are evenly distributed around an enlarged-diameter section of the riser. As the gas-solid system flows higher, it encounters another variable-diameter section where another group of nozzles are placed to achieve mix temperature control (MTC). MTC is achieved by injecting a selected liquid cut at

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low temperature to cool down both catalyst particles and gases. MTC can divide the riser into two regions: (i) an upstream zone with high mixing temperature of feedstock and regenerated catalysts, high catalyst to oil ratio, and a short contact time, and (ii) a downstream zone with

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lower temperature and milder reaction conditions. Because of the high temperature in the upstream zone, the amount of unvaporized feedstock can be lowered significantly and feed

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conversion can be promoted correspondingly. Due to the milder conditions in the downstream zone, higher LPG and propylene yield can be achieved whereas gasoline olefin can be further

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consumed so as to increase gasoline octane number. It is believed that MTC is an effective tool

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to obtain a high mixing temperature without raising the riser outlet temperature, hence avoiding excessive cracking reactions. [26] The upper part of the riser is a straight pipe with a constant

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diameter and 2 injection nozzles for slurry back wash oil. They are installed near the exit of the riser to terminate cracking reactions. Since the slurry wash back oil is heavier and has better absorption capacity on the catalyst surface than the rest of reactants, it can terminate cracking reactions and decrease riser temperature to avoid thermal cracking. It is notable that the geometry of the riser plays a crucial role in determining the adequate gas-solid regimes deemed

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necessary to achieve maximum conversion. When cracking reactions take place as the result of the feed coming into contact with regenerated catalysts, the resulting smaller molecules lead to a volume expansion in the gas phase. This expansion of vapor increases significantly both gas and catalyst velocities as they travel upward together along the FCC riser height. The expanding

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lighter vapor also vastly decreases the volume fraction of catalyst in the suspension and correspondingly reduces the catalyst to oil ratio (CTO). The changing diameter geometry should favor contact efficiency between catalysts and gaseous reactants and ensure an appropriate residence time for both phases within the FCC riser. [27]

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3. Hydrodynamics and Chemistry Governing Equations

A three dimensional reactive two-phase flow simulation was performed using the EulerianEulerian Two-Fluid model (TFM) to investigate the interplay between hydrodynamics and oil

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cracking chemistry within the simulated large-scale riser reactor of an industrial FCC unit. In the TFM approach, gaseous oil and solid catalyst phases are considered as mutually penetrating

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continua. Volume fractions with a sum of unity are introduced for both phases and they are

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assumed to be continuous functions of time and space. Conservation equations for continuity, momentum, species, and energy are solved for each phase, as shown in Table 1. Solid phase

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conservation equations are closed using constitutive equations derived from kinetic theory for granular flow (KTGF). The constitutive equations are summarized in Table 2. The network of

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cracking reactions is described using the 11 lump kinetic model [25] developed for a FCC unit for maximizing propylene. The use of such a kinetic model is consistent with the simulated riser operating conditions and production expectations. It is worth mentioning that feed oil was assumed to be completely vaporized at the inlet of feed injectors and hence no liquid phase is present at the start of the simulation. This is justified by the fact that for a 100 µm feed droplet,

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the time needed for complete vaporization varies from 0.3 to 30 ms in typical FCC units [28] which is far less than the feed oil residence time within typical industrial risers (3-4 s). Another reason is to be consistent with the fact that reaction parameters for the 11-lump kinetic model used herein and most other published lump kinetic models are derived with the hypothesis of

3.1 Conservation Equations Table 1. Governing equations for the gas-solid flow

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gaseous feed injection. [29] The schematic modeling framework is described in Figure 2.

Governing equations for Eulerian-Eulerian TFM

Continuity equation for

  g g      g  gVg   Wcoke t

(1)

  s s      s sVs   Wcoke t

(2)

  g  gVg      g  gVgVg  t   g P   g  g g    g   gs (Vg  Vs )

(3)

  s sVs      s sVsVs  t   s P   s  s g    s   gs (Vs  Vg )

(4)

 ( g  gY j )    ( g  gVgY j ) t    [ g  g DY j ]  W j

(5)

 ( s  sYcoke )    ( s  sVsYcoke ) t    [ s  s DYcoke ]  Wcoke

(6)

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Parameter

Momentum equation for

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gas and catalyst phase

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gas and catalyst phase

Component continuity equation

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for gas phase

Component continuity equation for coke in catalyst phase

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 ( g  g C pg Tg )    ( g  gVg C pg Tg ) t n

Energy equation for

   [Tg ]  W j Qrj  Qsg

(7)

j 1

gas and catalyst phase

(8)

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 ( s  s C psTs )    ( s  sVs C psTs )  Qsg t

In the current simulation work, the coke produced is considered to deposit on the solid catalyst

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phase. Accordingly, the catalyst deactivation model incorporated into the reaction model described below is COC (Coke on Catalyst) deactivation model and the mass fraction of coke in the catalyst phase is utilized consequently. With this in mind, there is a mass transfer between the gas and solid phases, which is equal to the production of coke, as shown in Eq. (1) and Eq. (2).

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For the catalyst phase, the reaction source term for coke is coupled to it since coke is deposited

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on catalyst particles (granular phase). So, the continuity equation for coke in the granular phase can be described as Eq. (6). In the gas energy equation, Eq. (7) accounts for the endothermicity

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of cracking reactions in gas phase which reads

n

W Q j 1

j

rj

and the heat transfers between the

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two phases which reads Qsg (J/m3·s). The “Gunn” model which is frequently used for Eulerian

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multiphase simulation involving a granular phase is chosen here to compute the Qsg . Qrj is the reaction heat (J/kg) for each reaction path.

3.2 Constitutive Equations To close the governing equations, appropriate closure laws for stress tensors  g , s and

momentum exchange coefficient  gs in momentum equations of both phases should be

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specified. In the closure of solid stress tensor  s , properties for granular phase such as granular pressure and granular viscosity are obtained using KTGF through the concept of granular temperature. Table 2. Constitutive equations for the KTGF KTGF correlations used in the CFD model

Granular temperature

3  [   s s s       s sVs s ] 2 t

equation

   Ps I   s  Vs    ks  s   s  gs





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Parameter

(9)

2kdil 6   [1  g0 s 1  e ]2  2 s 2  s d s 1  e  g 0   (10) 1  e  g0 5  

k s Diffusion coefficient

Collision

dissipation

1 75   s d s s 2 384

 

3 12(1  e2 ) g0  s s2s 2 ds 

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distribution

Shear viscosity

Collisional viscosity

g0 

1

1   s /  s ,max 

1/3

(11)

(12)

Ps   s s s 1  2 g0 s (1  e)

Solids pressure

function

kdil 

s

energy

Radial

where

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for granular energy

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1/2

(13)

,  s ,max  0.63

s  s ,col  s ,kin, gran  s , fric

(14)

(15)

where

s , fric  0

s ,col

4     s 2  s d s g0 1  e   s  5  

1

2

(16)

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10  s d s s 4 [1  g0 s 1  e ]2 96 1  e  g0 5

s ,kin, gran 

Bulk viscosity

s   s  s d s g0 1  e 

4 3

Stress tensor for gas phase Stress

tensor

for

s

(18)



2 3

 g   g  g {[Vg  (Vg )T ]  (Vg ) I }  s  [ g Ps   g s (Vs )]I 2  s s {[Vs  (Vs )T ]  (Vs ) I } 3

granular phase

 g  0.74 (21)

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 g  0.74

0.0214  0.576  4(  0.7463) 2  0.0044 (0.74   g  0.82) g  exchange  0.0038   0.0101  (0.82   g  0.97) (22) 4( g  0.7789) 2  0.0040  31.8295  32.8295 ( g  0.97) g  

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coefficient

(20)

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Interphase

(19)

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 3  g  s  g Vg  Vs   Cd ds 4  gs    s  g Vg  VS  s2  g  150  1.75 2   g ds ds  where

(17)

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Kinetic viscosity

Re 

 g  g d s Vg  Vs

(23)

g

0.687  24   1  0.15  Re   Cd   Re  0.44

Re  1000

(24)

Re  1000

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The solids stress tensor contains shear viscosity  s and bulk viscosity s arising from particle momentum exchange in the collision process. The expression for collisional viscosity s ,col and kinetic viscosity s ,kin, gran in shear viscosity  s are given by Gidaspow [30].The expression for

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bulk viscosity s in solids stress tensor is given by Lun et al. [31]. The momentum exchange between the two phases is based on the value of the interphase exchange coefficient  gs .Herein, we used the drag model based on the energy minimization multiscale (EMMS) approach proposed by Yang et al. [32]. This model captures the clustering

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behavior of the gas-solid systems as it was proved experimentally [32].

3.3 Reaction kinetic Model

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The selection of the kinetic model for the cracking reactions depends significantly on feedstock properties. The FCC riser reactor simulated herein treats a feedstock with very similar

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properties to the one that was used to derive the 11-lump kinetic model for maximizing propylene, [1, 25] especially in terms of Conradson carbon residue‟ and „hydrogen composition‟.

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Properties for the feedstock are listed in Table 3. As the main properties of the feed converted in

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the simulated riser match reasonably well that of the feedstock for the 11-lump model, it is logical to use the 11-lump kinetic model in the present work. All the lumped components in the

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11-lump kinetic model are listed in Table 4 and the reactions network is shown in Figure 3. Kinetic parameters as well as reaction heat for all the 41 reaction paths are summarized in Table 5.

The 11-lump kinetic model was incorporated into the 3D CFD model with the following assumptions: (1) all the reactions are irreversible; (2) cracking reactions of heavy oil and diesel

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oil are second-order reactions whilst other reactions are first-order reactions; (3) the coke-oncatalyst (COC) model based on the mass fraction of coke deposited on catalyst phase was chosen to establish the catalyst deactivation function; (4) catalyst deactivation which is caused by coke

Table 3. Feedstock properties Property

Feedstock used to derive 11-lump model

Density (293.15 K), (kg/m3)

0.905 2

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78.84(353.15 K)

Viscosity, mm2/s

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deposition is unselective.

38.83(373.15 K)

Conradson carbon residue, %

5.31

Carbon(C)

85.21

Hydrogen(H)

12.59

0.31

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Nitrogen(N)

0.23

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Sulphur(S)

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Elemental composition

Lump symbol

Boiling range

Heavy oil

HO

350-500℃

Diesel oil

DO

204-350℃

Olefin

G=

C5-204℃

Aromatics

GA

C5-204℃

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Lump

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Table 4. Lumped components of the 11-lump kinetic model [1]

Gasoline

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Saturates

G0

C5-204℃

Butane+propane

C3,40

C30+C40

Butylene

C4 =

C4 =

Propylene

C3 =

C3 =

DG=

C2 =

Dry gas Ethene

Ethane+Methane+H2 DG0 Coke

C20+C1+H2

CK

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LPG

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The reaction rate of each component (2nd order for heavy oil and diesel, and 1st order for other lumped component) can be expressed as below:

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dy  a  RCO  k  y dt

(26)

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dy  a  RCO  k  y 2 dt

1 1   (1  14.36CC ) 0.20 3.68 N 2.10 Ah 1 1 100 RCO 100 RCO

(27)

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a

In the above equations, a is the catalyst deactivation function that represents the level of

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catalyst activity, RCO is the catalyst to oil ratio, k is the reaction constant which is equal to

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k0e(  E / RT ) , k0 is the pre-exponential factor (for second order reaction its unit is y 1  s 1 whilst for first order reaction is s 1 ), E is the apparent activation energy ( kJ / mol ), y is mass fraction of the reactant, N and Ak are the mass fraction of nitrogen and (resin+asphaltene) in the feedstock which are 0.1% and 22.64% respectively, CC is the mass fraction of coke deposited on

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catalysts (During the coding of the 11-lump kinetic model, concentration of coke in the granular particle phase is utilized). Table 5. Kinetic parameters of the 11-lump kinetic mode [l1] H r (kJ / kg )

Reaction Path

1

HO → DO

601.20

59.14

180.27

2

HO → G=

1.16×104

79.48

914.80

3

HO → GA

1.58×106

117.17

795.52

4

HO → G0

470.83

63.74

720.97

5

HO → C3,40

0.43

23.67

1 348.07

6

HO → C4=

14.69

35.32

1 205.55

7

HO → C3=

8.29

30.44

1 690.13

8

HO → DG=

6.59

38.79

2 659.30

9

HO → DG0

2.12×105

114.75

3 822.30

10

HO → CK

28.49

11

DO → G=

12

DO → GA

13

DO → G0

14

DO → C3,40

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Ea (kJ / mol )

M

k0

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Path No.

0.00

2.68×103

75.32

734.53

2.04×105

108.27

615.24

63.15

540.69

2.98

31.61

1 167.80

DO → C4=

324.39

61.95

1 025.28

DO → C3=

542.64

65.35

1 509.86

ED

47.10

17

CE

PT

252.49

DO → DG=

1.27×104

97.64

2 479.03

18

DO → DG0

2.05×103

80.66

3 642.03

19

DO → CK

271.60

64.40

-180.27

20

G= → GA

128.29

76.48

-51.28

21

G= → C3,40

4.28

55.17

186.27

22

G= → C4=

61.12

58.69

125.00

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16

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G= → C3=

574.14

74.72

333.33

24

G= → DG=

89.57

68.12

750.00

25

G= → DG0

0.45

47.86

1 250.00

26

GA → C3,40

5.46

53.23

237.56

27

GA → C4=

48.64

59.38

176.28

28

GA → C3=

33.10

57.62

384.62

29

GA → DG=

6.12

48.32

801.28

30

GA → DG0

73.92

73.96

1 301.28

31

GA → CK

0.61

24.26

-342.01

32

G0 → GA

0.96

32.68

32.05

33

G0 → C3,40

0.69

32.60

269.61

34

G0 → C4=

167.09

66.27

208.33

35

G0 → C3=

105.06

63.51

416.67

36

G0 → DG=

324.79

73.04

833.33

37

G0 → DG0

2.13×103

94.81

1 333.33

38

C3,40 → DG=

2.03×105

117.70

563.73

39

C3,40 → DG0

565.03

74.11

1 063.73

40

C4= → C3=

1.18×103

90.51

208.33

41

C4= → DG0

80.14

1 125.00

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ED

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351.03

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23

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**Units are ( yi1  s 1 ) for the 2nd order reactions and s 1 for 1st order reactions.

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4. Numerical Model 4.1 Simulation System The geometry of the simulated large-scale FCC riser reactor is shown in Figure 1.

Regenerated catalysts are conveyed from an oblique pipe into the riser and then lifted by the prelifting steam injected at the riser bottom. Above the straight pipe section of one third of the riser

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height, there is a frustum cone-like diameter enlargement section where a series of feedstock nozzles are evenly distributed with a certain angle relative to the riser axis. A second conical section is used to enlarge the riser diameter with a group of MTC nozzles located at the top of this enlargement. At the upper section of the riser, there is a straight pipe section where slurry

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back nozzles are located at the end of it. Both MTC and slurry back nozzles are set with larger angles relative to the axial direction than that of feed injecting nozzles (slurry nozzle angle > MTC nozzle angle > feed nozzle angle). Even though the whole riser was simulated herein, special focus was given to the feed injection zone and the bottom conical enlargement section

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with feed nozzles since it is the region hosting most of the oil cracking process and it is characterized by a complicated interaction between the two phases. The total mesh number of the entire riser was about 865,000 with a greater grid density located in the diameter enlargement

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section where the flow pattern and reaction behavior were affected by the injected feed at a high speed. For grid convergence purpose, two coarser grids (180,000 and 375,000 cells) were

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originally tested and the results in terms of velocity, temperature, volume fraction profiles, and

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yield distribution are summarized in the Appendix.

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4.2 Boundary Conditions

The main boundary conditions are listed in Table 6. Feed oil was assumed to be injected at the

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vaporization temperature (Tevp= 460 C) similar to what was reported in the literature. [15,33,34] All the inlets were specified as “velocity inlet” whilst the riser outlet was set as “pressure outlet”. No-slip wall condition was set for gas phase while partial-slip condition was adopted for granular phase when specified the boundary condition for the „wall‟. Velocity for the gaseous feed oil, MTC feed, slurry back wash oil and catalysts were specified at the velocity inlets. Volume

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fraction of catalysts at the catalyst inlet, equal to 0.6, was calculated according to the flux of both regenerated catalysts and conveying steam in the oblique pipe.

4.3 Numerical Procedure The 3D geometry and mesh were created in the grid generation tool, GAMBIT 2.4.6 and then

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imported into the ANSYS Fluent 14.0 software. All the differential equations were solved using the finite volume method and discretized by a first order upwind differencing scheme. Cracking reaction terms and endothermic reaction heats were added as source terms to the species conservation equations and energy equation by means of user defined functions (UDFs) written

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in the C programming language. A small time step of 0.1 ms was used and a maximum of 100 iterations for each time step was chosen to achieve numerical convergence. To obtain a fully developed flow field inside the riser, the simulation time was chosen to be one order of

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magnitude longer than the average gas oil residence time inside the riser. Transient data were time-averaged over the last 10 s of the simulation. Simulations were run for 30 s on a High

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Performance Computer (HPC) populated with 64 computing nodes. Real simulation time was 3 months.

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Parameter

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Table 6. Boundary conditions

Prelift steam flux(kg/s)

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Prelift steam inlet temperature (K)

Value 0.82 633.15

Prelift steam inlet velocity (m/s)

0.33

Regenerated catalyst Flux (kg/s)

1 740

Regenerated catalyst inlet temperature (K)

973.15

Regenerated catalyst inlet volume fraction

0.60

Catalyst particle average diameter (µm)

70

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Regenerated catalyst inlet velocity (m/s)

0.94

Gaseous feed oil injection velocity (m/s)

55.60

Gaseous feed oil temperature(K)

733.15

MTC feed injection velocity (m/s)

59.43 693.15

Operating pressure (kPa)

223

5. Simulation Results and Discussion

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5.1 Hydrodynamic behavior

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MTC feed Temperature (K)

The velocity vector distribution of the gaseous oil phase on a cut-plane is shown in Figure 4. It can be seen that in the lower diameter enlargement section, the high speed injection of gaseous feed into the riser greatly influence the flow field in the whole riser. In this region, since the feed

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nozzles are located at the lateral wall, high speed of the jetting oil appears near the wall and then

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expands towards the inner center region gradually as the flow moves upward. This is also well depicted in Figure 5, which shows cross-section planes at different heights above the injection

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zone ( h is used here to represent the height above the feed oil or MTC injecting positions). Although the velocity magnitude for injected gas oil is high according to Figure 4, the radial

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velocity component is relatively quite low which contributes little to the radial momentum

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exchange. As a result, the radial distribution for gas phase velocity shows high inhomogeneity as depicted by the cross-section planes in Figure 5. Figure 6 shows that, in the straight pipe between the two conical sections, the gas velocity near the wall region is more uniform, but still much higher than that of the core region as a consequence of the fast feed injection at the wall. To investigate further the influence of larger injection angle (which is herein defined as the angle between the injection nozzle and the riser axis) on the two-phase mixing, it is important to

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look also at the radial velocity distribution above the MTC nozzles (whose angle with the riser axis is much larger than that of the lower feed atomizing nozzles) located at the upper conical enlargement section. Figure 7 shows the effects induced by the high velocity MTC nozzles on the radial velocity magnitude distribution. For the MTC nozzles, the injection speed is slightly

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higher and the injection angle is much larger than those of the feed nozzles. These two aspects, especially the larger angle of MTC nozzles, will result in a larger radial velocity component for gas oil and hence lead to a better momentum transfer in the radial direction as revealed by the results displayed in Figure 7. The latter shows also that the distance needed for the gas oil to

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have a homogeneous radial velocity magnitude distribution above the MTC nozzles is much shorter than that for the feed nozzles.

From the perspective of improving mixing of phases in the feed injection zone, which will

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further decrease the gradients of velocity, and benefit heat transfer as well as reaction behavior along the radial direction above the feed atomizing nozzles, the angle of the feed nozzles can be

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modified. According to several publications [14,16,35] on conventional FCC risers, feed injection angle has a more pronounced effect on the hydrodynamics of the feed mixing zone than

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injection position and injecting speed. It is reported that feed injection angle equal or wider than

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30° is desirable. For the present case, as the riser diameter is larger than those conventional risers, the current feed atomizing nozzle angle is not wide enough compared with the MTC

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nozzles to achieve lower gradients in the radial direction. This feed injection angle needs to be somewhat larger, for instance 45°, to give rise to a greater radial velocity component and hence a more homogeneous flow pattern. The latter often favors better heat transfer and chemical reactions. It is believed that this modification of feed injection angle would improve the operation conditions and economically benefit the production process.

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In Figure 7, the effect of the bottom high velocity of the injected feed still cannot be eliminated along the riser height resulting from the large flux of the feed and the high number of nozzles used, higher velocity near the wall region compared to the riser center still exists at the level of the MTC nozzles and above. At the upper enlargement section, the increase of residence time of

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gas oil because of the increase in riser diameter benefits the radial momentum transfer and gas velocity tends to be more homogeneous in the radial direction. Cracking reactions result in large molecules being cracked into smaller ones and thus lead to a volumetric expansion along the riser height. Consequently, the upper straight pipe with enlarged diameter offers a suitable region

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for the fluid to achieve a desirable fully developed flow pattern.

Figure 8 depicts the radial distribution of axial solid velocity, namely Z velocity component of solid phase at various heights. Based on this figure, one cannot observe the traditional core-

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annulus structure characterizing often the flow within riser reactors having feed oil injected at the bottom. This is due mainly to the mode of feed injection that consists of a group of evenly

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distributed feed nozzles located above the bottom prelifting zone. At this zone, the hydrodynamic behavior of catalysts is dramatically changed because of the high speed jetting

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gas. As shown in Figure 8, the maximum axial solid velocity appears in the near wall region

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rather than in the core region. Moreover, the effect of the fast injected gas oil gradually decreases as we move upward and inhomogeneity decreases accordingly. As mentioned above, in the upper

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straight pipe section, the fully developed flow pattern is achieved which decreases in general the difference between particle velocity in the vicinity of the wall and the riser center. Figures 9 and 10 show the volume fraction profile of catalysts in the whole riser and solid phase volume fraction distribution at different cross-section planes, respectively. As they encounter the high velocity jetting gas oil, catalyst particles near the nozzle inlets and in the extension of the nozzle

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direction are lifted up whilst others being trapped in the riser core. Accordingly, the volume fraction in the vicinity of the wall appears to be much lower than that in the central area. The lifted particles will encounter the enlargement section and disperse in the fluid zone, at the same time, efficient contact between catalyst and oil gas will be achieved which in turn benefit

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reacting performance, transport phenomena and volume expansion. In the upper region of the riser, solid volume fraction tends to be homogeneous in the radial section and decreases significantly because of the intensified cracking reactions as well as the improved transfer processes.

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5.2 Heat Transfer

Figure 11 details the temperature distribution of the gas phase. As discussed in the hydrodynamics section, the radial distribution for both phase velocities shows high

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inhomogeneity as depicted in Figure 5 and Figure 8, and as heat transferring phenomena and reaction performance are closely interrelated with hydrodynamics, inhomogeneous temperature

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distribution can consistently be observed in Figure 11, in which low temperature appears near the wall whereas higher temperature characterizes the riser center. As high speed jetting feed flows

PT

along the near wall region, endothermic cracking reactions mainly occur near the wall where the

CE

low temperature region is consequently observed. A wider feed injection angle would allow more gaseous feed flow into the core region and then benefit the temperature distribution in the

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radial direction.

5.3 Species Concentration and Reaction Characteristics Predicted product compositions at the outlet of the simulated primary riser and the process

values of the product yields are listed in Table 7. It can be seen that the errors between the simulated data and the process data are acceptable. In the dual riser process, the primary riser is

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demonstrated to play the leading role in converting the feedstock with a conversion of 78.99%, and the yield of liquid products approaching 65.34% after going through the primary reactor. The predicted propylene composition by the model at the outlet of the primary riser is slightly higher than the process value (12.45% against 10.20%) and is deemed satisfactory. This slight

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discrepancy can be attributed to the fact that the 11-lump kinetic model used in the simulation of the FCC riser for maximizing propylene favors those reactions generating propylene in order to characterize the propylene promotion in the cracking process. Also, the prediction of olefin in gasoline is relatively high if compared to the reported industrial data. [7] This may be linked to

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the injection angle of the feedstock nozzles. Indeed, for the current relatively small feed injection angle, the resulting axial velocity component would be high whereas the radial velocity would be quite low. This often leads to a short residence time in the whole riser height and inhomogeneous

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velocity, temperature, and species distribution in the radial direction. The above often reduces the secondary reactions responsible for converting olefins in gasoline to propylene. Larger feed

ED

injection angle would induce more gaseous feed oil to run near the riser center and reduce the axial velocity component. The latter will eventually increase the gas residence time inside the

CE

PT

riser and thus more gasoline olefin will be converted into propylene.

Table 7. Comparison of different products‟ mass fraction at the primary riser outlet

Process data

(wt%)

(wt%)

Absolute

Diesel oil

11.95

11.41

0.54

4.73

Gasoline

24.93

26.40

-1.47

-5.57

LPG

24.01

27.53

-3.52

-12.79

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Simulation results

Term

Errors Relative (%)

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8.54

5.20

3.34

64.23

Coke

5.86

8.45

-2.59

-30.65

Unconverted feedstock

24.71

21.01

3.70

17.60

Conversion

75.29

78.99

-3.70

-4.68

100

100

0.00

0.00

Propylene

12.90

10.20

2.70

26.47

Olefin in gasoline

12.03

-

-

Yield of liquid products

66.75

65.34

1.41

-

2.16

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Sum

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Dry gas+loss

Figure 12 and Figure 13 depict the mass fraction distribution and variation rate of mass fraction for the four main products, namely diesel oil, gasoline, LPG and dry gas. The variation

y k , ( yk  s 1 ): t

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rate of species mass fraction is defined as

ED

yk yk ,t  yk ,t t  t t

(28)

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yk is the variation of mass fraction for species k , yk ,t is the mass fraction for k at the

CE

current time step, yk ,t t is the mass fraction for k at the last time step, t is a time step size, s.

y k was stored in a User Defined Memory (UDM) that was chosen to be a custom t

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The value of

field function to allow time averaging during the simulation process and thus obtain meaningful predictions.

y k can represent the net reaction rate for a species considering both generating and t

consuming rates. Furthermore, Figure 13 depicts the main generation zone of the four main products. For diesel oil, its reaction rate is lower than the other products. This is attributed to the

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fact that catalytic cracking reaction is a parallel and sequential scheme where the intermediate species of diesel will further be converted to lighter products. Consequently, the yield of diesel oil is generally low in the main body of the riser except for the vicinity of feed oil nozzles where diesel is significantly produced as revealed in Figure 12. From Figure 13, reaction rate of

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gasoline is high in the whole riser whilst those of LPG and dry gas are high mainly in the upper section of the riser apart from the feed oil injecting zone. This is expected since the production of LPG and dry gas mainly rely on the secondary reactions which occur in the upper section of the riser. Although the generation zone of the dry gas is similar to that of LPG, the absolute value of

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dry gas is far more less than LPG.

It is easy to recognize differences among reaction rates for olefin, aromatics and saturates in gasoline product from Figure 14. It also explains the selectivity differences of the three lumped

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components. Cracking reactions give a priority to the generation of olefin as revealed in Figure 14. It is well known that even though olefin in gasoline reduces gasoline octane number, it

ED

effectively benefits the yield of propylene. As cracking reaction progresses, olefin content first increases and then falls down as a result of propylene conversion. In this large-scale FCC riser,

PT

the reaction rate of gasoline olefin starts to decrease only at the riser top. This can be traced back

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to the feed injection angle which results in a high axial velocity and in turn lower the residence time, consequently, no enough time is left for the conversion of olefin in gasoline to propylene. It

AC

can also be revealed from the pre-exponential factor for propylene generation reaction in the 11lump kinetic model that olefin in gasoline has a faster converting rate to propylene than to other light olefins. For the sake of optimization, feed injection angle is of vital importance in terms of residence time necessary for secondary reactions to take place and further produce propylene. Figure 15 shows a higher selectivity for propylene compared to the other two lumped

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components. As mentioned above, olefins in gasoline notably benefit generation of propylene in addition to the conversion of butylene to propylene in the sequential cracking process.

6. Conclusions

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A 3D numerical model was built to study the interplay of hydrodynamics, heat transfer, and catalytic cracking chemistry within a novel large-scale FCC riser reactor. The CFD model was based on the Eulerian-Eulerian two-fluid model to simulate the hydrodynamics of gas and

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catalyst particle phases and 11-lump kinetic reaction model to account for the cracking chemistry. A novel treatment for the coke component was conducted by incorporating coke into the secondary granular phase which is more realistic since carbon deposition occurs on catalyst phase. Simulation results regarding the distribution of the different yields were compared to

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process data and good agreement was achieved. Detailed description of both hydrodynamics and reaction behaviors highlighted the effects nozzle injecting angle has on flow and thermal patterns

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at the feed injection zone. Wider angle for feed nozzles was recommended since they would

PT

induce higher radial velocities which in turn benefit mixing in radial direction, better heat transferring condition, and desirable reaction performance. In future investigations, quantitative

CE

evaluation of the dynamic cracking process should be a focus. Instead of solely concentrating on the final product yield distribution at the outlet of the riser and assessing the cracking

AC

performance inside the riser qualitatively in accordance with the riser geometry and operation conditions, a new unified and quantitatively evaluative method that can incorporate all impact factors and give an indicator of the riser performance is urgently needed. By this new method, and using a finer grid, cracking performance of separate riser regions or different risers with various conditions can be evaluated and straightforwardly compared.

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AUTHOR INFORMATION Corresponding Author

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Fax: +971-260-75408 (A.S.B.)

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E-mail: [email protected] (A.S.B)

ACKNOWLEDGMENT

Financial support for this project provided by TAKREER (Abu Dhabi Oil Refining Company)

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operating company of the Abu Dhabi National Oil Company (ADNOC) and National Basic

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Research Program of China (2012CB215006) is gratefully acknowledged.

ABBREVIATIONS drag coefficient

AC

Cd

CE

Nomenclature

Cp

specific heat, J/kg·K

D

Coefficient of diffusion, m2/s

ds

averaged granular diameter, μm

e

coefficient of restitution

g

gravity, m/s2

30

radial distribution function

I

unit tensor, Pa

P

static pressure, Pa

Qr , j

reaction heat of component j, J/kg

Qsg

heat transfer between two phases, J/m3·s

Re

Reynolds number

t

time, s

T

temperature, K

V

velocity, m/s

Wcoke

reaction rate of coke in the secondary phase, kg/m3·s

Wj

reaction rate of component j, kg/m3·s

Ycoke

mass fraction of coke

Yj

mass fraction of component j

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g0

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Greek Letters volume fraction



density, kg/m3

s

granular temperature, m2/s2



stress



interphase exchange coefficient



drag force modify coefficient



shear viscosity, Pa·s



bulk viscosity, Pa·s

AC

CE

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

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

energy exchange coefficient between two phases

g

gas phase

s

solid phase

i, j

component name

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Subscripts

APPENDIX

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A. Comparison of different products’ mass fraction at the primary riser outlet for the coarse grid (180,000 cells)

Process data

(wt%)

(wt%)

Diesel oil

12.49

Gasoline

24.81

LPG

ED

Errors

Absolute

Relative (%)

11.41

1.08

8.65

26.40

-1.59

-6.41

24.19

27.53

-3.34

-13.81

8.05

5.20

2.85

35.40

PT

M

Simulation results

5.76

8.45

-2.69

-46.70

24.69

21.01

3.68

14.09

75.31

78.99

-3.68

-4.89

AC

Term

100

100

0.00

0.00

Propylene

12.45

10.20

2.25

18.07

Olefin in gasoline

11.51

-

-

Yield of liquid products

67.26

65.34

1.92

Dry gas+loss Coke

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Unconverted feedstock Conversion Sum

2.85

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B. Comparison of different products’ mass fraction at the primary riser outlet for the medium grid (375,000 cells) Simulation results

Process data

(wt%)

(wt%)

Absolute

Diesel oil

12.45

11.41

1.04

Gasoline

24.71

26.40

-1.69

-6.40

LPG

24.21

27.53

-3.32

-12.06

Dry gas+loss

8.04

5.20

2.84

54.61

Coke

5.66

8.45

-2.79

-33.13

Unconverted feedstock

24.93

21.01

3.92

17.09

Conversion

75.07

78.99

-3.92

-4.96

100

100

0.00

0.00

10.20

2.31

22.65

12.51

Olefin in gasoline

11.69

Yield of liquid products

67.03

ED

Propylene

Errors Relative (%) 9.11

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Sum

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Term

-

-

65.34

1.69

2.59

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The results of the coarse and medium grids are compared in terms of velocity, temperature and volume fraction profiles in different regions of the riser in order to assess the effect of the grid

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on the results. The average differences on all of the above between the two grids were 17%, 14%, 11% respectively. When we looked only at the lower part of the riser (only cells below the

AC

second set of nozzles), these differences were 11%, 10%, 9% respectively. For the upper part (only cells above the second set of nozzles), these differences were 21%, 15%, 11% respectively. Note that we do not have any process data on the velocity, temperature and void fraction in any part of the riser. These differences in the hydrodynamics did not have noticeable effects on the yield distribution as predicted by the model using both coarse and fine grids. The two tables above show the

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error in matching the process data using both coarse and medium grid. The agreement is good for most of the yield for both grids Regarding the use of the finer grid (865,000 cells), the average differences on velocity, temperature, and void fraction between the medium grid and the fine grid (375k and 865k) were 16%, 14%, 12% respectively. When we looked only at the lower part of the riser (only cells

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below the second set of nozzles), these differences were 11%, 8%, 8% respectively. For the upper part (only cells above the second set of nozzles), these differences were 20%, 15%, 12% respectively.

The matching between the model prediction using the fine mesh (865k Cells) and real data in

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terms of yield distribution is not much different from the one of the medium mesh (375k Cells).

REFERENCES

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[1] J. Q. Gan, H. Zhao, S. B. Abdallah, C. H. Yang, H. H. Shan, Numerical Simulation of Hydrodynamics and Cracking Reactions in the Feed Mixing Zone of a Multiregime Gas–Solid

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Riser Reactor, Ind. Eng. Chem. Res. 50 (2011) 11511-11520.

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[2] Y. H. Xu, J. S. Zhang, J. Long, A Modified FCC Process MIP for Maximizing Iso-paraffins

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in Cracked Naphtha, Pet. Process. Petrochem, 32 (2001) 1-5. [3] Y. M. Chen, Recent advances in FCC technology, In Fluidization and Fluid Particle

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Systems, Proceeding of Annual Meeting of the American Institute of Chemical Engineers, Austin, United States, Nov 7-12, 2004; Elsevier: Amsterdam, 2006; 2-8. [4] G. F. Ramos, J. M. Fusco, E. F. Sandes, M. Einsfeldt, M. J. Bampi, A New Way to Deal

with Thermal Balance in Fluid Catalytic Cracking Process, In Proceeding of 16th World Petroleum Congress, Rio de Janeiro, Brazil, 2002.

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[5] C. Y. Li, Z. W. Xu, G. H. Jiang, H. Z. Ding, G. H. Wang, Y. J. Wang, C. H. Yang, Commercial test of two-stage riser catalytic cracking of heavy oil for maximizing propylene yield, Petrochem. Technol. & Appl. 26 (2008) 436-441.

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[6] C. Y. Li, C. H. Yang, H. H. Shan, Maximizing propylene yield by two-stage riser catalytic cracking of heavy oil, Ind. Eng. Chem. Res. 46 (2007) 4914-4920.

[7] L. Y. Wang, B. L. Yang, G. L. Wang, H. T. Tang, Z. B. Li, J. L. Wei, New FCC process minimizes gasoline olefin, increases propylene, Oil Gas J. 101 (2003) 52-58.

AN US

[8] J. Sun, H. H. Shan, C. W. Li, X. Chen, Study on the catalytic pyrolysis of coker gas oil and coker naphtha upgrading based on TMP technology, Pet. Process. Petrochem. 42 (2011) 46-51. [9] J. Q. Gan, C. H. Yang, C. Y. Li, H. Zhao, Y. B. Liu, X. Luo, Gas–solid flow patterns in a

M

novel multi-regime riser, Chem. Eng. J. 178(2011) 297-305.

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[10] Z. Y. Liu, Z. D. Zhang, X. H. Gao, H. T. Zhang, Industrial application of LCC-300 catalyst for enhancing propylene yield in two-stage riser catalytic pyrolysis of heavy oil, Pet.

PT

Process. Petrochem. 42 (2011) 47-50.

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[11] X. H. Li, C. Y. Li, J. F. Zhang, C. H. Yang, H. H. Shan, Effects of temperature and catalyst to oil weight ratio on the catalytic conversion of heavy oil to propylene using ZSM-5 and

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USY catalysts, J. Nat. Gas Chem. 16 (2007) 92-99. [12] J. Chang, F. D. Meng, L. Y. Wang, H. F. Yang, Theory and Practice of Low Temperature

Contact and High Ratio of Catalyst to Oil in FCC Process, Acta Petrolei Sinica Sect. Pet. Process, 27 (2011) 367-375.

35

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[13] K. N. Theologos, N. C. Markatos, Advanced modeling of fluid catalytic cracking riser‐ type reactors, AIChE J. 39 (1993) 1007-1017. [14] J. Chang, F. D. Meng, L. Y. Wang, K. Zhang, H. G. Chen, Y. P. Yang, CFD investigation

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of hydrodynamics, heat transfer and cracking reaction in a heavy oil riser with bottom airlift loop mixer, Chem. Eng. Sci. 78 (2012) 128-143.

[15] J. S. Gao, C. M. Xu, S. X. Lin, G. H. Yang, Y. C. Guo, Advanced model for turbulent

AN US

gas–solid flow and reaction in FCC riser reactors, AIChE J. 45 (1999) 1095-1113.

[16] J. Chang, W. Z. Cai, K. Zhang, F. D. Meng, L. Y. Wang, Y. P. Yang, Computational investigation of the hydrodynamics, heat transfer and kinetic reaction in an FCC gasoline riser, Chem. Eng. Sci. 111 (2014) 170-179.

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[17] J. Li, Y. P. Fan, C. X. Lu, Z. H. Luo, Numerical Simulation of Influence of Feed Injection

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on Hydrodynamic Behavior and Catalytic Cracking Reactions in a FCC Riser under Reactive Conditions, Ind. Eng. Chem. Res. 52 (2013) 11084-11098.

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[18] B. Chalermsinsuwan, P. Kuchonthara, P. Piumsomboon, CFD modeling of tapered

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circulating fluidized bed reactor risers: Hydrodynamic descriptions and chemical reaction responses, Chem. Eng. Process. Sect. Process Intensification, 49 (2010) 1144-1160.

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[19] X. L. Zhu, C. H. Yang, C. Y. Li, Y. B. Liu, L. Wang, T. Li, Q. Geng, Comparative study

of gas–solids flow patterns inside novel multi-regime riser and conventional riser, Chem. Eng. J. (2013) 215-216, 188-201.

36

ACCEPTED MANUSCRIPT

[20] S. Chen, Y. P. Fan, Z. H. Yan, W. Wang, C. X. Lu, CFD simulation of gas-solid twophase flow and mixing in a FCC riser with feedstock injection, Powder Technology, 287(2016) 29-42.

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[21] K. Xiong, C.X. Lu, Z.F. Wang, X.H. Gao, Quantitative correlations of cracking performance with physiochemical properties of FCC catalysts by a novel lump kinetic modelling method, Fuel, 161(2015) 113-119.

[22] G. C. Lopes, L. M. Rosa, M. Mori, J. R. Nunhez, W. P. Martignoni, CFD study of

AN US

industrial FCC risers: the effect of configurations on hydrodynamics and reactions, Int. J. Chem. Eng. (2012) 1-16.

[23] H. C. Alvarez-Castro, E. M. Matos, M. Mori, W. Martignoni, R. Ocone, Analysis of

M

process variables via CFD to evaluate the performance of a FCC riser, Int. J. Chem. Eng. (2015) 1-13.

ED

[24] J. Li, F. Z. Xiao, Z. H. Luo, A CFD modeling of the gas-solid two-phase flow in an FCC

PT

riser under the electrostatic conditions, Asia-Pacific J. Chem, Eng. 9(2014) 645-655. [25] J. H. Guo, Primary study of the lumped kinetic model for heavy oil cracking into

CE

propylene by two-stage-riser technology. M. S. Thesis, China University of Petroleum (East

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China), Dongying, China, 2007. [26] C.F. Dean, J. L. Ross, Improvements to increase fuel capacity and quality, Pet. Technol.

(2002) 29-38.

[27] Q. H. Liu, W. Sun, G. L. Niu, C. H. Yang, Study on particle flow characteristics in new structure FCC riser, Pet. Refinery Eng. 37 (2007) 32-36.

37

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[28] J.S. Buchanan, Analysis of heating and vaporization of feed droplets in fluidized catalytic cracking risers, Ind. Eng. Chem. Res. 33 (1994) 3104-3111. [29] H. Zhang, Numerical Simulation of the Gas-solid Two-phase Flow in TMP Diameter-

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changing Reactors. M. S. Thesis, China University of Petroleum (East China), Qingdao, China, 2008.

[30] D. Gidaspow, R. Bezburuah, J. Ding, Hydrodynamics of Circulating Fluidized Beds, Kinetic Theory Approach, In Proceedings of the 7th Engineering Foundation Conference on

AN US

Fluidization, Gold Coast, Australia, May 3-8, 1992; United States, 1992, 75-82.

[31] C. K. K. Lun, S. B. Savage, D. J. Jeffrey, N. Chepurniy, Kinetic theories for granular flow: inelastic particles in couette flow and slightly inelastic particles in a general flowfield, J.

M

Fluid Mech. 140(1984) 223-256.

[32] N. Yang, W. Wang, W. Ge, J. H. Li, Choosing structure-dependent drag coefficient in

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modeling gas-solid two-phase flow, China Particuol. 1 (2003) 38-41.

PT

[33] C. Zhu, Y. Jun, R. Patel, D. W. Wang, T. C. Ho, Interactions of flow and reaction in fluid

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catalytic cracking risers. AIChE J. 57 (2011) 3122-3131. [34] C. N. Wu, Y. Cheng, Y. Jin, Understanding riser and downer based fluid catalytic

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cracking processes by a comprehensive two-dimensional reactor model, Ind. Eng. Chem. Res. 48 (2008) 12-26.

[35] J. Li, Z. H. Luo, X. Y. Lan, C. M. Xu, J. S. Gao, Numerical simulation of the turbulent gas–solid flow and reaction in a polydisperse FCC riser reactor, Powder Technol. 237 (2013) 569-580.

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