Short Contact Time Review Papepr

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I NTERNATIONAL J OURNAL OF C HEMICAL R EACTOR E NGINEERING Volume 3 2005 Review R1 A Review of Short Residence Time Cracking Processes Craig Hulet * Cedric Briens Franco Berruti Edward W. Chan ** * The University of Western Ontario, [email protected] University of Western Ontario, [email protected] University of Western Ontario, [email protected] ** Syncrude Canada Limited, [email protected] ISSN 1542-6580 Copyright c 2005 by the authors. All rights reserved. No part of this publication may be reproduced, stored in a retrieval system, or transmitted, in any form or by any means, electronic, mechanical, photocopying, recording, or otherwise, without the prior written permission of the publisher, bepress, which has been given certain exclusive rights by the author.

Transcript of Short Contact Time Review Papepr

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INTERNATIONAL JOURNAL OF CHEMICAL

REACTOR ENGINEERING

Volume 3 2005 Review R1

A Review of Short Residence TimeCracking Processes

Craig Hulet∗ Cedric Briens†

Franco Berruti‡ Edward W. Chan∗∗

∗The University of Western Ontario, [email protected]†University of Western Ontario, [email protected]‡University of Western Ontario, [email protected]∗∗Syncrude Canada Limited, [email protected]

ISSN 1542-6580Copyright c©2005 by the authors.

All rights reserved. No part of this publication may be reproduced, stored in a retrieval system,or transmitted, in any form or by any means, electronic, mechanical, photocopying, recording, orotherwise, without the prior written permission of the publisher, bepress, which has been givencertain exclusive rights by the author.

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A Review of Short Residence Time CrackingProcesses

Craig Hulet, Cedric Briens, Franco Berruti, and Edward W. Chan

Abstract

This review examines the key features and configurations of short residencetime cracking processes from a diverse range of industries that have been de-veloped over the past 25 years. These industries include: bitumen or heavy oilupgrading, biomass pyrolysis, olefin production, catalytic cracking, and coal gasi-fication. Characterization of the gas, liquid, and solid products and feedstock isprovided wherever possible. In addition, a description of the source and mech-anism of heat transfer, and how the feedstock is brought into contact with andseparated from – this source is also given.

There is a strong economic incentive for considering short residence time crack-ing processes. Not only do such processes increase the yields of the more valuableliquid and gaseous products, but more compact designs would also decrease cap-ital costs. Careful control of the vapour residence times appears to be crucial inorder to prevent secondary cracking and yet allow for maximum cracking of thefeedstock. Rapid and thorough mixing of the feedstock with the heat source, notjust creating a uniform dispersion, is also a key design aspect to consider. Finally,rapid and complete separation must also be carefully considered; again, to helpcontrol product residence time and avoid secondary cracking but also from a heatbalance point of view.

KEYWORDS: catalytic, coal, contacting, cracking, biomass, bitumen, heavy oil,olefins, pyrolysis, separation, short residence time, thermal conversion, upgrading

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1 INTRODUCTION

Considerable attention is being directed towards fast pyrolysis processes in a wide variety of industries. Many terms are used in the literature to describe such systems, including ‘ultra-pyrolysis’, ‘short residence time’, ‘ultra-rapid pyrolysis’, and ‘quick contacting pyrolysis’. However, the basic principle remains the same: heating the given feedstock rapidly in the absence of air until it vaporizes or cracks. In this manner, solid or heavy liquid carbonaceous feedstocks may be transformed into more valuable liquid and gaseous components with the production of less solid material in a shorter period of time. These products are typically used for fuels or chemical feedstock.

Fast pyrolysis systems are distinguished from the traditional pyrolysis methods not only on the basis of residence time, but also by the temperature, and the rate of heating, to which the feedstock is subjected. Although individual definitions will vary, a fast pyrolysis system will generally possess the following characteristics:

• a heat transfer rate on the order of 1000°C/s (1832°F/s) to rapidly heat the feedstock to the required pyrolysis temperature;

• a residence time, measured from the period of initial contact between the feedstock and the heat source to the time of product quench, on the order of a few milliseconds and generally not exceeding 3 s;

• and rapid separation from the heat source and quenching of the product vapours in order to minimize secondary cracking and the production of less valuable solid carbonaceous materials.

To this end, many different configurations have been developed (or are under development) to meet these objectives and provide the proper conditions required to maximize the production of the more valuable gas and/or liquid products.

This review describes and summarizes the key features and results obtained for the major fast pyrolysis systems that have been developed over the past 25 years. Particular attention has been paid to the source and mechanism of heat transfer and how the feedstock is brought into contact with this source. Other considerations include the reactor configuration, flow regime, and the techniques used for separation and quench of the product vapours. When available, the feedstock characteristics, and the typical results obtained are given and their relationship to the operating severity characterized.

The fast pyrolysis systems that have been included in this discussion have been categorized according to use into the following five categories: heavy oil upgrading, catalytic cracking, olefin production, biomass pyrolysis, and coal pyrolysis. However, this categorization is not definitive and some processes, with or without any modifications, may actually be placed in one or more categories. In addition, only those processes meeting the criteria discussed above were included, with one relevant exception.

2 HEAVY OIL UPGRADING FAST PYROLYSIS PROCESSES

2.1 Internally Circulating Fluidized Bed

2.1.1 Introduction

The spouted bed reactor provides a unique method for contacting fluids and solids. The basic premise of the spouted bed consists of introducing a single centralized jet vertically from the bottom of the reactor vessel into a fluidized bed of particles. As the fluid jet rises, it creates a distinct path through the bed and entrains some of the surrounding solid particles. As it clears the surface the jet swells outwards and loses momentum and is no longer capable of maintaining the solid particles in suspension. In a standard spouted bed the solids are in continuous contact with the fluid jet, hence the jet must be of sufficient force to support the bed of surrounding material without collapsing, thereby limiting the maximum bed height (or maximum spoutable bed height). The maximum spoutable bed height also appears to be a function of temperature (Milne, 1993). Furthermore, gas from the central jet will

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“leak” out into the surrounding annular region of bed solids. These characteristics would make such a reactor unsuitable for short contact time cracking.

In light of these concerns, the internally circulating fluidized bed (ICFB) was developed at The University of Calgary (refer to Figure 2.1) for use in studying short contact time pyrolysis of heavy oil. The pilot design includes a central axially aligned draft tube placed within the fluidized bed of solids just above the inlet of a central feed jet and extending above the bed surface. Solids enter the riser by passing through a series of orifices located at the riser bottom (Milne et al., 1994). This feature eliminates any concern with spout stability and maximum spoutable bed height that is common with conventional spouted beds as mentioned above and eliminates the effect of temperature on the maximum spoutable bed height (Milne, 1993). The annular fluidized bed of solids maintains a net positive pressure difference between the annular region and the riser thereby minimizing the amount of spouting gas bypassing the riser and entering the annular region of the bed. In this manner mass fluxes on the order of 1000 kg/m2 (205 lb/ft2) are achievable (Milne et al., 1994). The feed used in the studies by Milne and his colleagues was Amoco Ipiatik heavy oil.

2.1.2 Process Description

Figure 2.2 is a flow diagram for the high temperature pyrolysis process described in the study by Milne (1993). Overall, the reactor is 2.1 m (6.9 ft) in height with the diameter ranging from 0.15 m (0.49 ft) at the base to 0.20 m (0.66 ft) at the top. The riser dimensions include a diameter of 0.035 m (0.11 ft) and an overall height of 1.1 m (3.6 ft). Entering the reactor are fluidization lines for the annular region, the riser gas, the heavy oil feed, and the feed atomization gas. The fluidization and riser gases consisted of either steam or air.

Fluidization gasRiser gas

Riser assembly

Solids disengaging “T”

Gas baffle

Product gas

Figure 2.1: ICFB reactor (after Milne et al., 1994).

Feed

Auxiliary gas Riser gas

Atomization gas

Furnace

ICFB Reactor

Quench waterNitrogen

Primary separator

Product accumulation drum Condensate

trap

Quench chamber

Condensers

Figure 2.2: Flow diagram for the ICFB pyrolysis process (after Milne, 1993).

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The bed contents – either a catalytic or inert solid such as sand, depending on the desired product range – are heated by means of an external clamshell furnace. The furnace used was capable of achieving temperatures of 1200°C (2192°F) giving temperatures in the riser from 600 to 950°C (1112 to 1742°F) (Lovett et al., 1997). The preheated heavy oil, once sprayed onto the hot sand particles, quickly reaches the required pyrolysis temperature in the riser. In earlier designs, the solids were separated from the products using a “T” section at the top of the riser (see Figure 2.2). Exiting the reactor the hot product gases are immediately quenched using water atomized with nitrogen gas. The primary separator serves to collect any entrained solids and eventually any liquid product formed. The product stream terminates after passing through two final condensers and a pair of collectors.

2.1.3 Operating Conditions

The typical operating conditions for the reactor consisted of a feed oil flow rate to the reactor system of 1.0 bpd, a maximum residence time of 1000 ms, and a bed temperature of 800°C (1472°F). The reactor residence time is assumed to be accurate to within +/- 10%. The approximate droplet size used was between 1 to 10 µm with the spray nozzle located just underneath riser orifices in order to prevent coke deposits from forming on any internal surfaces. The gas temperature exiting the quench chamber was typically 300°C (572°F).

2.1.4 Product Characterization and Material Balances

It is stated by Milne (1993) that the performance of the oil feed and product collection systems was such that they impeded on the accuracy of the mass balance calculations. Other factors cited that may have affected the material balance calculations included problems cleaning the primary condenser and the formation of a hydrocarbon vapour observed to exit the system with the exhaust gas.

Table 2.1 below shows the final results of the analysis of the product emulsion for all four runs that were conducted. As may be observed from Table 2.1, the final product densities for all four runs are above 1000 kg/m3 (62.4 lb/ft3). The conclusion drawn from this observation was that the oil produced was highly aromatic. Table 2.2 summarizes the final product analyses that were conducted.

Table 2.1: Product emulsion analysis (after Milne, 1993). Case 1 2 3 4 Reactor temperature, °C (°F) 650 (1202) 658 (1216) 745 (1373) 724 (1335) Residence time, ms 604 454 548 668 Oil fed, kg (lb) 2.3 (5.1) 5.6 (12.3) 3.1 (6.8) 6.1 (13.4) Product emulsion, kg (lb) 3.1 (6.8) 5.7 (12.6) 1.5 (3.3) 7.8 (17.2) Oil, wt % 32.2 36.9 59.2 20.3 Water, wt % 65.8 59.7 35.9 77.0 Solids, wt % 2.0 3.4 4.9 2.7 Total estimated coke, kg (lb) 0.16 (0.36) 0.30 (0.65) 0.17 (0.38) 0.31 (0.68) Product oil yield, wt % 42.8 37.9 29.1 26.2 Coke yield, wt % 7.0 5.3 5.1 5.6

The coke contained, on average, 3.3-wt % ash. The coke content remaining on the sand particles ranged from 0.2- to 0.4-wt %. Overall, the inside of the reactor was relatively coke-free with minor deposits of coke found at the gas outlet, instrumentation ports, and at the quench spray nozzle. In general, the mass balance was poor since the mass of hydrocarbon product collected represented only 40% of that fed to the system. In a subsequent analysis, the product oil from run 4 was combined with the corresponding condensate. The recombined oil had a viscosity of 19.6, 9.8, and 3.8 mPa-s at 25, 50, and 100°C (77, 122, and 212°F), respectively. The density was 1060 kg/m3 (66.2lb/ft3) at 25°C (77°F). Table 2.3 and Table 2.4 show the results of the distillation analysis that was performed on the recombined oil sample and how it compares to the feed and uncombined oil sample in terms of product fractions.

2.1.5 Waste Plastic Ultrapyrolysis

The pilot-plant scale ICFB reactor discussed above has also been applied in several studies of upgrading of plastic wastes (e.g. Milne et al., 1999; Sodero et al., 1996). The overall goal of such research is the development of an

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Table 2.2: Final product properties (after Milne, 1993). Heavy oil feed 1 2 3 4 C, wt % 82.1 81.2 81.1 85.1 82.4 H, wt % 10.3 8.01 6.35 5.56 5.49 N, wt % 0.51 0.61 0.64 0.66 0.56 S, wt % 4.75 4.68 4.61 5.04 4.46 residue, wt % 2.31 5.55 7.26 3.65 7.13 carbon residue, wt % 10.8 20.9 27.1 32.1 23.9 asphaltenes, wt % 15.6 24.7 34.9 45.9 34.4 maltenes, wt % 84.4 75.3 65.1 54.1 65.6 viscosity, cP at 25°C (77°F) 9540 93.7 69.2 186 158 at 50°C (122°F) 1341 40.4 33.2 56.6 54.3 at 100°C (212°F) 86.1 12.1 11.3 11.7 12.6 density (kg/m3) at 25°C (77°F) 983 1050 1080 1140 1120 Simdist IBP, °C (°F) 221 (430) 216 (421) 217 (423) 218 (424) 210 (410) IBP – 300°C (572°F), % 7.8 15.6 12.4 9.2 13.3 300 – 400°C (752°F), % 16.6 25.3 17.9 15.0 18.7 400 – 525°C (977°F), % 21.1 20.4 15.2 11.5 12.2 + 525°C (977°F), % 54.4 38.7 54.6 64.2 55.8 Gas produced N/A Analysis, wt % Methane 36.7 19.3 23.4 Ethane 7.2 5.1 7.7 Ethylene 17.7 33.3 33.0 Propylene 22.0 21.5 21.0 Butane 10.5 10.4 5.8 Butadiene 5.9 10.3 9.0

efficient and economic means of recycling plastics for the recovery of valuable chemical products such as liquid fuel or even the base monomer itself.

Table 2.3: Distillation of recombined oil using ASTM D1160 (after Milne, 1993).

Fraction Boiling range wt % Naphtha 86 (IBP) – 204°C

(187 – 400°F) 18.9

Light gas oil 204 – 343°C (400 – 650°F)

21.8

Heavy gas oil 343 – 514°C (650 – 957°F)

34.3

Total distillables

86 – 204°C (187 – 400°F)

75.0

Residues +514°C (+ 957°F) 22.6 Loss – 2.4

Table 2.4: Product fraction comparison (after Milne, 1993).

Feed Initial oil

Recombined oil

Naphtha 0.0 0.0 19.4 Light gas oil 14.7 21.8 22.3 Heavy gas oil 29.2 21.7 35.1 Total distillables

43.9 43.5 76.8

Residue 56.1 56.5 23.2

Feed rates for the studies were 2.88 kg/hr (6.35 lb/hr) of powdered low-density polyethylene (LDPE) and 2.30x104 – 2.48x104 kg/s (5.08x104 – 5.48x104 lb/s) of sand, which acted as the heat carrier. The gas superficial velocity through the riser ranged from 7.4 to 8.4 m/s (24.3 to 27.6 ft/s). Temperatures within the riser were varied from 780 to 860°C (1436 to 1580°F) while the residence time was between 400 to 600 ms. Table 2.5 summarizes some of the key results of this study. In general, gas yields exceed 90-wt %; with a total olefin yields of around 75-wt % making them the primary product (Milne et al., 1999). Included in Table 2.5 is a comparison to other similar

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Table 2.5: Summary of product yields (wt% of feed) from plastic waste pyrolysis using ICFB (after Milne et al., 1999).

ICFB reactor (Sodero et al., 1996).

ICFB reactor (Lovett., 1997).

WFBB (Scott et al., (1990)).

Temperature, °C (°F) 797 (1467)

825 (1517)

780 (1436) 805 (1481)

730 (1346) 790 (1454)

Residence time, ms 400 400 600 600 500 500 Methane 10.9 10.7 22.3 23.3 4.6 6.4 Ethylene 18.7 22.9 28.2 32.3 19.4 31.1 Propylene 13.4 13.4 17.8 17 12 12.8 Other alkanes 1.8 1.8 1.3 1.8 3 3.5 C4 42.6 39 18.4 15.6 13.1 7.7 C5 + 5.4 5 0 0 6.3 0.7 Total olefins 74.7 75.3 64.4 64.9 44.5 51.6 Total gas yield 92.9 92.9 88 90 58.4 62.2 Condensate ~5 ~5 ~8 ~8 31.2 32.1 Char ~2 ~2 ~4 ~2 2.1 0.2 Total 99.9 99.9 99.9 90 91.7 94.5

studies by Lovett et al. (1997) and Scott (1990) using the ICFB and WFBB processes (see Section 5.2.2), respectively.

Another study by Lovett et al., (1997) involved mixing 5-wt % waste plastic with 95-wt % heavy oil in the ICFB reactor. The heavy oil used was from Cold Lake, Alberta. The residence time was held constant at 600 ms. The gas riser mass flux varied from 2.30x104 to 2.45x104 kg/hr (5.08x104 to 5.40x104 lb/hr), which corresponds to a superficial velocity of 7.4 to 8.4 m/s (24.3 to 27.6 ft/s) as mentioned above.

The major products included methane, ethylene, propylene, butene, and butadiene. Overall gas yields increase from 60 to 83% with increasing temperature, which was varied between 750 and 830°C (1382 and 1526°F). It was found that by adding 5-wt % LDPE increased the overall gas yields by about 5-wt % but did not significantly effect the ethane, ethylene, or propylene yields, which were consistently between 50- and 55-wt %. By contrast, both the butene and butadiene yields increased by about 5-wt % over the given temperature range. Finally, the methane yields decreased with increasing temperature as much as 10 – 15-wt % at high temperatures (Lovett et al., 1997).

These later studies describe the use of new patented technique that allowed for the separation of the solids from the gaseous vapours in under 20 ms with an efficiency of 99.5% or greater (Nielsen et al., 1997). Figure 2.3 shows an illustration of the device, which takes the form of a semi-circular inverted U. Figure 2.3 (a) illustrates an inclined surface used to create a venturi effect within the separator, immediately after this ramp is a divided chamber. Gas escapes

(a) (b)

Figure 2.3: Riser terminator developed for ICFB (a) side view (b) perspective view (after Nielsen et al., 1997).

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through a series of slotted openings while the solids continue along the upper curved surface and leaves through the downwardly directed opening. The gas is collected via the through the discharge outlet located to the one side

2.2 Thermovortactor

2.2.1 Introduction

Ultrapyrolysis is a fast pyrolysis process developed at The University of Western Ontario. Built at a mini-pilot plant scale, the system was initially designed for the pyrolysis of biomass (Graham et al., 1986), however the process was recently applied to bitumen as well (Vogiatzis, 1994). For this study the heat carrier used was nitrogen gas although a solid heat carrier, such as sand, was used for other feedstocks (Shen et al., 1995).

2.2.2 Process Description

Figure 2.4 shows the simplified flow plan for the system. Nitrogen, heated to 100°C (212°F) above the reactor temperature is injected into the upper portion of the reactor where it comes into contact with the liquid feed, which itself has been heated to approximately 200°C (392°F). The vortical reactor promotes rapid mixing which allows rapid heat transfer to occur so that the temperature required for pyrolysis is reached in about 20 ms. The components then enter the cylindrical isothermal reaction zone, which is approximately 1 m (3.3 ft) in length and defines the fixed residence time of the system. The mixture then enters a quench chamber and is sprayed with cold nitrogen. Finally, the liquid and any entrained liquid droplets are separated using a cyclonic condenser, electrostatic precipitator, and filter before the gases are collected for sampling.

Q u e n c h c h a m b e r

R e a c t o r c o l u m n

V o r t i c a l r e a c t o r

P r e h e a t e d f e e d

P r e h e a t e d N 2

N 2 q u e n c h

C y c l o n i c c o n d e n s e r

E l e c t r o s t a t i c p r e c i p i t a t o r

T o f i l t e r g a s c o l l e c t i o n b a g s

Figure 2.4: Thermovortactor flow plan (after Vogiatzis, 1994).

2.2.3 Reactor Operating Conditions

The reactor operating conditions are given in Table 2.6. The reactor residence times were varied by the use of cylindrical stainless steel inserts and adjustments in the nitrogen gas flow rates.

Table 2.6: Thermovortactor operating conditions (after Vogiazis, 1994).

Operating temperature 500 – 900°C (932 – 1652°F) Reactor mixing times 10 – 20 ms Reactor residence times 70 – 500 ms Quenching time 20 ms Feed rates 0.20 kg/hr (0.44 lb/hr)

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2.2.4 Results

Liquid quality analyses were only conducted on a few select runs in order to determine the amounts of nitrogen, sulphur, nickel, and vanadium contaminants as well as the liquid viscosity. In addition, these analyses were limited to the 500 – 750°C (932 – 1382°F) temperature range due to a marked decrease in the H/C ratio above 750°C (1382°F). Liquid yields as high as 90% were obtained when the reactor temperature was less than 650°C (1202°F). In general, an increase in the reaction temperature resulted in a decrease in the liquid yield. An increase in the reaction temperature also leads to a decrease in the H/C ratio, which otherwise remained relatively close to that of the feed. On average a decrease from 6390 cSt to 100 cSt at 40°C (104°F) was observed between the feed and the liquid product (for a reaction temperature of less than 750°C (1382°F)). Some typical results from the liquid analyses are given in Table 2.7.

Table 2.7: Liquid quality for the Thermovortactor (after Vogiatzis, 1994). Run T Residence

time, Viscosity H/C

(liq) C

(wt %) S

(wt %) N

(wt %) Ni Va

(#) °C (°F) ms cSt at 40°C (104°F)

wt/wt wt % liquid wppm wppm

3 500 (932) 274 - 0.118 84.16 3.63 - 15.7 17.8

47 550 (1022) 153 - 0.129 84.32 3.73 - 24.0 22.3

31 555 (1031) 456 - - - 3.61 - - -

1 610 (1130) 241 110 - - 3.73 0.285 - -

44 650 (1202) 159 - 0.131 83.29 3.96 - 33.4 34.8

4 655 (1211) 271 34 - - 3.54 0.298 - -

2 690 (1274) 226 46 - - 3.68 0.302 - -

33 700 (1292) 451 - 0.114 83.38 4.04 - 24.5 24.8

5 750 (1382) 255 158 - - 3.99 0.449 - -

34 750 (1382) 422 - 0.089 85.51 4.34 - 26.4 25.8

25 900 (1652) 73 - - - 5.33 - - -

As for the gas product, it was observed to be primarily olefinic (with ethylene yields of up to 18% under the most severe conditions) with a H/C ratio of between 0.21 and 0.28. The H/C ratio was seen to decrease as the reaction temperature increases.

Elemental levels of sulphur and nitrogen found in the liquid product remained relatively high for a reaction temperature of less than 750°C (1382°F), however the levels were lower than that of the feed. The sulphur was primarily removed as H2S. The initial values of nitrogen and sulphur in the feed were 0.45-wt % and 4.4-wt %, respectively. By comparison, the levels of nickel and vanadium were significantly lower than in the feed (originally 77 wppm nickel and 190 wppm vanadium).

It was observed that between 70 and 80% of the nickel and vanadium were removed, as well as a relatively small amount of the sulphur. The H/C ratio was between 0.024 and 0.037 and did not seem to be affected by the reaction temperature. Table 2.8 summarizes some key solid properties for selected runs. Finally, Table 2.9 summarizes the product yields for some selected runs.

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Table 2.8: Solid properties for the Thermovortactor (after Vogiatzis, 1994). Run T C/H (liq) C H S (#) °C (°F) wt/wt wt % solid wt % solid wt % solid 3 500 (932) 0.024 58.0 15.7 26.2

47 550 (1022) 0.037 56.6 15.8 27.6 44 650 (1202) 0.026 66.2 15.3 18.5 33 700 (1292) 0.024 78.3 16.3 5.4 34 750 (1382) 0.024 79.1 16.7 4.2

Table 2.9: Product yields for the Thermovortactor (wt %) (after Vogiatzis, 1994). Run T Residence Gas yield Liquid yield

°C (°F) time, ms H2 H2S C1 – C4 Total C5+ total Solid yields

18 600 (1112) 238 ND 0.400 1.730 2.13 1.84 90.87 7.00 11 610 (1130) 330 ND 0.363 2.587 2.95 2.80 90.05 7.00 51 700 (1292) 176 0.026 0.749 8.435 5.62 5.62 83.79 10.59 50 700 (1292) 236 0.065 1.016 13.319 8.66 8.66 78.60 12.74 49 700 (1292) 286 0.064 1.175 15.301 9.10 9.10 76.46 14.44 48 700 (1292) 340 0.079 1.224 16.287 10.32 10.32 75.41 14.27 36 710 (1310) 137 ND 0.750 12.720 12.70 12.70 79.53 7.77 57 750 (1382) 182 0.112 1.222 18.226 11.20 11.20 73.44 15.36 56 750 (1382) 222 0.148 1.541 22.731 13.10 13.10 68.58 18.32 5 750 (1382) 255 ND 0.982 23.348 10.82 10.82 68.67 20.51

55 750 (1382) 293 0.171 1.610 24.579 13.70 13.70 66.64 19.66 54 800 (1472) 188 0.223 1.381 24.906 10.13 10.13 66.49 23.38 53 800 (1472) 208 0.279 1.701 29.210 10.80 10.80 61.81 27.39 38 800 (1472) 373 0.490 1.700 36.740 12.40 12.40 54.07 33.53

2.3 Thermal Regenerative Cracking and Quick Contact Cracking

2.3.1 Introduction

Quick contact (QC) cracking is a development of Stone and Webster Engineering Corp (SWEC) and Gulf Oil Chemicals. It is assumed here to be an evolutionary extension of the Thermal Regenerative Cracking (TRC) process also developed jointly between SWEC and Gulf Oil Chemicals (Ellis et al., 1982; Gartside and Ellis, 1983; Gartside and Norton, 2002). The primary difference is that for the TRC process the heat carrier particles are inert, while for the QC process they are catalytic.

A large-scale pilot unit, referred to as the field-test unit (FTU), was constructed in 1981 capable of handling 1270 kg/hr (1.4 tons/hr) of hydrocarbon feed and 45360 kg/hr (50 tons/hr) of solids (Gartside, 1989). The process is reputed to be relatively insensitive to coking making it suitable for heavy feedstocks. Initially designed for olefins production, a recent patent now refers to the process as a low residence time cracking process for liquid fuels (Johnson et al., 1999) indicating that the process is flexible. Depending upon the desired product and exact process configuration the inert solid might be replaced with catalytic solids instead. The system is capable of handling a wide range of feedstocks including vacuum residue bottoms and any heavy hydrocarbon feed containing sulphur, heavy metal contaminants, and coke precursors (Johnson et al., 1999). Byproduct recycle has also been found to be a viable option (Hu, 1982).

The QC process consists of a high velocity downflow reactor where active or inert particulate solids are used as a heat carrier. The reported residence time of the process is 200 ms from the initial feed injection to the time of product quench (Gartside, 1989). The downflow reaction scheme was selected for its close approximation to plug flow and low pressure drop. The key remaining process equipment include a patented separator design, solid-feed

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mixer and standpipe level control system, effluent quench section, and an entrained bed heater regeneration zone (Johnson et al., 1999).

2.3.2 Process Description

Figure 2.5 shows a simplified process diagram for the most recent configuration of the QC system. To begin, the solids enter the tubular reactor via the secondary regenerator and entrained bed heater (EBH). To facilitate feeding they are fluidized using steam having by this point been heated to between 704 and 982°C (1300 and 1800°F). The reactor pressure is between 101 to 2514 kPa (15 to 365 psia) and the residence time between 0.02 and 0.5 s. The preferred solid to liquid ratio for the feed is 5 to 15, which will provide the desired cracking temperature of between 538 and 649°C (1000 to 1200°F). The steam to hydrocarbon ratio is 0.2 to 0.4. Next, the solid particles and reactions products enter the separator which returns the former to the EBH after first passing through the stripper collector while the latter is sent for further processing as required.

2.3.2.1 Solids-feed mixer

Figure 2.6 shows a detailed cross-section of the patented solids-feed mixer. The solids enter from the vertical conduits shown while the liquid feed enters from the surrounding toroidal feed line through angled openings. The openings are preferably convergently beveled so that the spray pattern forms a cone shape enveloping the entering solids while at the same time is directed away from the walls in order to minimize erosion. A mixing plug serves to enhance this rapid and intimate mixing by reducing the

Separator

Steam

Air

Entrained bed heater

Flue or fuel gas

Secondary regenerator

Air Air

HCReactor feeder Reactor

Figure 2.5: The quick contact (QC) flow plan

(after Johnson et al., 1999).

Hydrocarbon feed

Fluidization steam

Central core plug

Figure 2.6: The QC solids-feed mixer (after Johnson et al., 1999).

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flow area thereby forming discrete mixing zones (Johnson et. al., 1999). The reactor feeder is lined with an inner core of ceramic material.

2.3.2.2 Separator

Detailed views of the solid separation system are given in Figure 2.7 and Figure 2.8. The solids separation system is reportedly to have a 99.99% removal efficiency with over 95% of the solids removed within 30 ms (Gartside, 1989; Johnson et al., 1999). At the entrance to the separator lies a 90° right angle and so as the solids impinge against the opposing wall a mass of solids accumulate upstream of the weir. The bed of solids protects the inner wall of the separator. The gas acts to sweep the upper layer of particles towards the solids outlet then exits the separator orientated at a 180° angle to the initial inlet. The solids flow is by gravity, preferably in a downward direction. Further details of the separator design are presented by Johnson and in US patent 4,288,235 (as referenced in Johnson et al., 1999).

Collected solids

Solids outlet

Gas outletInlet

Weir

Figure 2.7: The QC separator (after Johnson et al., 1999). Figure 2.8: Separator side view (after Johnson et al., 1999).

2.3.3 Results

No documented evidence as to the actual performance of the QC system has been found in the published literature as of yet. The earlier publications state that this system was developed as a superior replacement to tubular reactors because this system could handle a wider range of feedstocks, was more flexible in the fuel used for heating, and allowed cracking to occur under more severe conditions. The patent from Johnson et al. (1999) states that the composition of the cracked products is similar to that obtained from a conventional FCC unit, but with certain benefits. These benefits, derived from the shorter residence time, include a reduction in coke yields and light gas yields while the yields of gasoline, diesel oil, and fuel oil products increase.

Table 2.10 shows a typical set of results used to illustrate US patent 5,976,355 (Johnson et al., 1999). The operating values listed are for the QC process and an unspecified “conventional” process for the treatment of an Arabian Atmospheric Tower Bottoms (ATB) feed. The results

Table 2.10: Typical operating conditions for treating an Arabian ATB feed (after Johnson et al., 1999).

Operating conditions QC process Conventional process Reactor temperature 566°C (1050°F) 950°F (510°C) Residence time 0.20 s 2.0 s Regenerated catalyst temperature 732°C (1350 °F ) 704°C (1300°F) Catalyst to feed weight ratio 9 8 Pressure 20 psig – Temperature of hydrocarbon feed 260°C (500°F) 93°C (200°F)

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Table 2.11: Typical yields for the case outlined in Table 2.10 (after Johnson et al., 1999).

QC process Conventional process C2 and lighter 1.7 3.5 C4 – gasoline 49.8 44.6 HCO (heavy cycle oil) 5.7 8.7 Coke 6.7 13.3 Conversion, vol. % 85% 77%

obtained using the QC system are compared to those obtained using the conventional process in Table 2.11.

2.4 Curie point pyrolyser

2.4.1 Introduction

The Curie point Pyrolysis Microreactor system was developed as part of a research program into the kinetics of short residence time hydrocarbon pyrolysis conducted at The University of Calgary. The system was first applied to studying the kinetics for the pyrolysis of propane (Rastogi et al., 1998) and then n-hexadecane (Fairburn at al., 1990). Tan (1993) then employed the system while conducting a kinetic study on the pyrolysis of Cold Lake and Peace River heavy oils. More recently studies on the upgrading of plastic wastes have been conducted using this process (Lovett et al., 1997; Sodero et al., 1996). Finally, Gray et al. (2003) have developed a model for devolatilization and cracking of Athabasca bitumen residues. Advantages of the batch micro-reactor system include a high degree of reproducibility and accuracy. Also, each experiment is performed in a relatively short time span that would allow many conditions to be rapidly examined along with any replicates.

The system is based upon the principle of high frequency induction heating – in this case by placing a ferromagnetic material in an induction coil and rapidly alternating the polarity of the applied current. The material will rapidly (e.g. tens of milliseconds) reach a point where the material loses its ferromagnetism – at the so-called Curie point. There is a unique Curie point for each ferromagnetic material of a given composition and wires are available with Curie point temperatures ranging from 300 – 1000°C (572 – 1832°F) (Milne, 1993).

2.4.2 Process Description

Figure 2.9 shows a diagram of the Curie point Pyrolysis system as used by Rastogi (1988). To begin, a piece of ferromagnetic wire is coated with approximately 15µg of the hydrocarbon sample and vacuum-sealed in a glass tube. The dimensions of the glass capsule are approximately 0.012 m (0.47 in) in length by 0.0007 m (0.028 in) in diameter. The wire itself is approximately 0.0004 m (0.016 in) in diameter and for the study by Tan (1990) was composed of a cobalt-nickel-iron alloy.

Microswitch Signal to stop timer Hammer activating solenoid

Hammer

Pyrolysis chamber

Microreactor

Heater

Inductive current coil

Sensor

Power supply

Heater supply

Carrier gas

Injection block

Gas chromatograph

Figure 2.9: The Curie point pyrolysis microreactor

system (after Rastogi, 1998).

During an experiment, the flow of carrier gas is directed into the reaction chamber. A delay timer then activates simultaneously both the reactor and a counter-timer. The residence time for the pyrolysis reaction is set using a hammer controlled by another delay timer, which breaks the glass vial and also stops the counter-timer. In this fashion all the response times for all components are taken into account and the time recorded is the actual residence time, not just the set residence time. The carrier gas, cold helium in this case, acts to quench the reaction before sweeping the product gases into a chromatography column for analysis. Finally, both the wire and glass fragments are recovered for analysis of any residue.

2.4.3 Results

Figure 2.10 shows some typical results from the study by Tan (1990) for ethylene yields obtained during the pyrolysis of Cold Lake heavy oil at temperatures ranging from 800 to

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R esidence tim e (m s)

200 400 600 800 1000 1200

Eth

ylen

e yi

eld

(wt%

of f

eed)

2

4

6

8

10

12

14

16

18

20

1000 deg.C .900 deg.C .800 deg.C .

Figure 2.10: Ethylene yields from the pyrolysis of Cold Lake heavy oil (after Tan, 1990).

1000°C (1472 to 1832°F). The residence times were varied from 250 to 1000 ms.

As concerns the study by Sodero et al. (1990) temperatures were varied between 800 and 900°C (1472 and 1652°F) with residences times ranging from 350 to 2500 ms while pyrolysing polyethylene. At 800°C (1472°F) and high residence times the relatively low gas yields (32%) were observed, while at 900°C (1652°F), the gas yields increased to 95-wt % for a residence time of 1000 ms. In general, increased reaction time gave increased yields of methane and ethane. Maximum yields for ethylene, propylene, and 1,3-butadiene were observed at residence time of 750 ms and 800°C (1472°F). In other words, maximum olefin production rates were observed at high temperatures and low residence times (Sodero et al., 1996).

Lovett et al. (1997) used this microreactor system for studying the pyrolysis of polystyrene at temperatures ranging from 800 to 965°C (1472 to 1769°F). The reactor residence times were maintained between 500 and 1000 ms. The goal of the study was to determine the conditions for maximizing the production of styrene and benzene. Once again, the optimal reactor production rates were observed at higher temperatures and lower residence times. For instance at 965°C (1769°F) and 500 ms the combined styrene and benzene yields were over 95%.

2.5 Cyclone Reactor Concepts

2.5.1 Introduction

One example of the use of cyclone reactor concept for the pyrolysis of bitumen has been created at the laboratory scale at University of Utah. For this study various tar sands were tested including samples from Sunnyside, Whiterocks, PR Spring, the Tar Sand Triangle, and Circle Cliffs. Bitumen saturation levels of these tar sands range from 8- to 14-wt %. The tar sand properties are given in Table 2.12 for two of the sands studied.

Once again, the most relevant operating parameters were found to be the reactor temperature and the residence time. Generally, it was discovered that at constant residence times the liquid yield decreased with

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increasing temperature and increased with decreasing residence time when the temperature was held constant (Lin et al., 1987).

2.5.2 Process Description

For the experiments the reactor temperature was approximately 550°C (1022°F) with an inlet gas temperature of 1003°C (1837°F), under ambient pressure. No express mention is made as to whether the cyclone is heated directly or just indirectly by the heated gases and preheated feed particles. The sand particles were fed either at ambient temperatures or preheated to 200°C (392°F). The experimental flow plan is given in Figure 2.11. The production rate was on the order of 3000 bpd with tar sand feed rates of 1.9x105 to 3.4x105 kg/hr (210 to 370 tons/hr). No detailed information is presented concerning the quenching, separation, and collection of the final products. The coked sand particles are sent to a cyclonic regenerator where the particles are preheated prior to reentering the reactor.

Table 2.12: Tar sand properties (after Lin et al., 1987). PR Springs Whiterocks Bitumen saturation 14.1 8.0 API gravity 7.8 10.3 Viscosity, cP 160 29.250 CCR, wt % 14.0 13.0 Ash content, wt % 3.3 0.8 Simulated distillation volatility, wt % 31.9 22.1 IBP - 400°F, wt % 1.3 0.9 400 – 650°F, wt % 5.1 3.3 650 - 1000°F, wt % 25.6 18.8 1000°F +, wt % 68.1 77.2 Elemental Analysis, wt % C 84.7 85.0 H 11.3 11.4 N 1.3 1.3 S 0.5 0.4 O 1.8 1.6 Ni 110 84 V 10 < 200 Atomic H/C ratio 1.61 1.61 Molecular weight, g/mol 702 NA Pour point, °C (°F) 104 (220) –

Sized bitumen impregnated sandstone 298 K

Cool spent sand < 367 K

1367 KClean spent sand

823 KSpent coked sand

Preheated air

1276 K

Preheated bitumen-impregnated sandstone 298/473 K

Fuel gas

Bitumen derived hydrocarbon liquid

Combustor off gas

Gas/liquid separator

Hydrocarbon vapor + gases 823 K

Pyrolysis cyclone reactor 823 K

Combustion cyclone reactor 1367 K

Figure 2.11: University of Utah process diagram (after Lin et al., 1987).

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2.5.3 Results

Typical results for the pyrolysis of the two tar sands listed in Table 2.12 are given below in Figure 2.12. It should be noted that the residence times shown in Figure 2.12 are estimated values of the gas residence time, and are therefore the upper maximum limit value for the solid residence time. However, these values are at least two orders of magnitude higher than those reported in similar systems outlined in Section 5.5. Overall, estimated yields for liquids are on the order of 60 – 80-wt %, 10 – 20-wt % for C1 to C4 gases, and 10 – 20-wt % as solid residue (Lin et al., 1987). Conversion was dependent to a large extent on residence time to a large extent due to the diameter of particle used in the study. For instance, a 4 m (13 ft) cyclone could achieve 90% conversion to liquid products when the particle diameter was around 3x10-4 m (0.012 in) but not for particles 0.001 m (0.039 in) in diameter.

To= 298 KTs= 823 KParameter: sand partic le diameterPR Spring Rainbow I tar sand deposit

100 20 30 40 50 60

20

40

60

80

100

Time (s)

Bitu

men

con

versi

on (%

)

0.025 cm

0.12 cm

0.2 cm

0.4 cm

To= 298 KTs= 823 KParameter: sand partic le diameterWhiterock tar sand deposit

100 20 30 40 50 60

20

40

60

80

100

Time (s)

Bitu

men

con

versi

on (%

)

0.4 cm

0.2 cm

0.12 cm

0.025 cm

(a) (b)

Figure 2.12: Typical results for the pyrolysis of the two tar sands shown in Table 2.12 (after Lin et al., 1987).

2.6 US Department of Energy

Figure 2.13: DOE tar sand pyrolysis system using a cyclone reactor

(after Westhoff and Harak, 1989).

Another cyclone reactor process for the pyrolysis of tar sands has been patented by the United States Department of Energy. This reactor system was developed for the rapid pyrolysis of tar sands of varying bitumen content of between 6 and 14-wt % using what appears to be a traditionally configured cyclone. Similar to the process described in Section 2.5 the heat required to initiate the pyrolysis reaction is supplied by heated sand particles recycled between the cyclone reactor and a cyclone shaped fluidized bed burner. A flow plan for this process is shown in Figure 2.13. The feed is introduced starting from the upper left where it immediately passes through a heat exchanger. The gaseous products exit the top of the cyclone and pass through a series of heat exchangers before entering a standard demister. These same heat exchangers are used to

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preheat a portion of the product dry gas that is recycled back to the burner along with a stream of fresh air. These streams are further heated by the accumulated ash exiting from the bottom of the burner.

Particle sizes used ranged up to 0.013 cm (0.5 in) in diameter. The process was originally developed in order to avoid problems inherent with earlier designs, in particular the inability to handle tar sands with a wide range of bitumen content, and mixing of combustion gases from the burner with product gases from the cyclone (Harak, 1982; Westhoff and Harak, 1989). At this period, the process economics were such that tar sands of 6.0-wt % bitumen would require all of the product gases and some of the liquid oil produced to supply the required energy. By comparison, a tar sand with 9.5-wt % would require less than half of the process gas, while at 14.14-wt % bitumen the process would require only a portion of the char that is produced (Westhoff and Harak, 1989). Table 2.13 presents a set of typical results for three different sets of pyrolysis runs using tar sands with bitumen contents of 6.0-, 9.5-, and 14.14-wt %. The tar sands used were from the Athabasca region of Alberta, Canada.

Table 2.13: Representative results from the DOE pyrolysis process (after Westhoff and Harak, 1989).

Tar sand properties Bitumen content, wt % 9.5 6.0 14.14 Feed rate, kg/hr (lb/hr) 282 (621) 446 (983) 189 (417) Potential oil yield, bbl/d 3000 3000 3000 Potential gas yield, kg/hr (lb/hr) 5 (12) 5 (12) 5 (12) Potential char yield, kg/hr (lb/hr) 3 (6) 3 (6) 3 (6) Feed streams x103, kg/hr (lb/hr) Tar sands 282 (621) ( – ) ( – ) Air 71 (157) ( – ) ( – ) Water 96 (212) ( – ) ( – ) Product streams x103, kg/hr (lb/hr) ( – ) ( – ) Product gas 3 (6) ( – ) ( – ) Liquid oil 19 (42) ( – ) ( – ) Sand (ash) 255 (562) ( – ) ( – ) Water 55 (120) ( – ) ( – ) Combustion products 76 (168) ( – ) ( – ) Steam 42 (92) ( – ) ( – ) Process results Char burned x103, kg/hr (lb/hr) 3 (6) 3 (6) 3 (6) Oil burned x103, kg/hr (lb/hr) 0 1 (2) 0 Gas burned x103, kg/hr (lb/hr) 2 (5) 5 (12) 0 Oil produced, bpd 3000 2831 3000 Gas produced x103, kg/hr (lb/hr) 1 (3) 0 5 (12) H2S free gas produced, m3/hr (ft3/h) 2400

(83639) 0 4360

(154007) Discharged sand (ash) x103, kg/hr (lb/hr) 254 (562) 416 (924) 358 (358)

2.7 Lurgi-Ruhrgas Process

The Lurhi-Ruhrgas process (LR process for short) was developed by the Lurgi Co., in Germany during the 1950’s. The basic process, as illustrated in Figure 2.14, was developed for retorting finely crushed oil shale in order to produce gaseous and liquid compounds suitable for use as fuels. Such a process was also referred to as a sandcracker since the heat carrier used was often grains of sand, but in practice any small sized particle such as coke particles or the recovered oil shale solids would do.

The hydrocarbon feedstock is preheated to between 346 and 400°C (655 and 750°F) using a superheater and injected into a sealed screw-type conveyor where it is mixed with the hot particulate solids. The reactor itself operates at about 704 – 843°C (1300 – 1550°F). The reactor residence time is approximately 0.3 to 0.5 s. Upon exiting the conveyor reactor the solids enter a surge bin while the vapors are withdrawn, passed through a series of cyclones then sent to the product recovery section. A portion of the solids is regenerated using a riser-type regenerator. The regenerated solids are collected in another surge bin and then gravity fed back into the reactor.

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The largest process unit built was at Dormagen, Germany in 1958 and had a production capacity of 1.4- to 1.8x107 kg/yr (3- to 4x107 lb/yr) of ethylene using a naphtha feed (Hu, 1982). Hu (1982) mentions that they were problems reported with the refractory lining in early testing. A typical set of cracking results is given in Table 2.14.

2.8 Fortum Oil & Gas Oy’s

The NexCC process developed by Fortum Oil & Gas Oy’s (formerly Neste Oy) was originally designed to process high viscosity gas oil (b.p. 320 – 400°C or 608 – 752°F). The end products obtained from the NexCC catalytic cracking process are olefins for gasoline components, alkylation, and etherification (Hiltunen et al., 2002).

The NexCC reactor is a circulating fluidized bed reactor that may be divided into a series of concentric annular regions that are axially aligned and vertically orientated. Referring to Figure 2.15, the NexCC reactor may be divided into three main regions: the product separation and withdrawal section, the mixing and reaction section, and finally the regenerator section. The product separation and withdrawal section includes the inner most cylindrical tube through which the gaseous products exit. It also includes the two annular chambers surrounding the product withdrawal line, one large and one small, which act as transfer channels for the solids recovered by the cyclone. The direction of flow within these three regions is from top to bottom. The multiport cyclone that separates the catalytic solids from the gaseous products exiting the reactor is located just above the product withdrawal line. The reactor and mixing section is located in the annular region that surrounds the product separation and withdrawal section. Here the direction of flow is from the bottom to the top with the injection of the hydrocarbon feed occurring at the very bottom as indicated in Figure 2.15 through vertically aligned feed nozzles.

Flue gas waste

Waste

Gaseous and liquid produc ts

Air and fuel (if required)

Oil shale

Solids to waste

Figure 2.14: The process flow diagram for the LR-process

(after Hendrickson, 1975).

Table 2.14: Cracking yields for the Lurgi-Ruhrgas process (after Hu, 1982).

Feedstock Iraqi Crude Cracking temperature, °C (°F) 1350 1390 Yield, wt. % Ethylene 19.6 23.1 Propylene 12.6 12.8 Butadiene 3.4 3.7 n & I Butadiene 4.7 2.0 C1 – C4 paraffins 13.8 14.4 Gases 3.4 2.4 Gasoline up to 220°C (428°F) 17.0 16.5 Oil above 220°C (428°F) 20.5 22.1 Carbon on sand and loses 5.0 3.0

As discussed in the previous paragraph, the solids that are separated by the inner cyclone pass down through two transfer channels. The outer and smaller of the two chambers may be used to return solids to the reactor if desired. The inner and larger of the two communicates between the inner reactor tube and the outer regenerator section via a control valve. The solids enter the riser regenerator found in the outermost annular region. The transport gas oxidizes the solids prior to entering another multiport cyclone. Both multiport cyclones have louvered inlet that create a centrifugal force on the solids. The separated solids then flow downward through two annular regions just to the inside of the regenerator. A portion of these solids may be returned to the regenerator while the majority reenters the reactor. The channel communicating back into the reactor is filled with a dense phase of solids in order to impede the flow of gases from the reactor into the regenerator.

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The reactor portion is generally maintained at a temperature between 650 to 1000°C (1202 to 1832°F) for thermal cracking and 520 to 650°C (968 to 1202°F) for catalytic cracking. The corresponding residence times are 0.2 – 0.5 s for thermal cracking and 0.5 – 5 s for catalytic cracking (Allan and Blaser, 1989). The NexCC offers a very compact design in comparison to a conventional riser due in part to the inline cyclone separators. Ruottu et al. (2000) also cite a decrease in the distance required for lateral mixing as compared to a conventional reactor with the same cross-sectional area. Table 2.15 summarizes a set of representative results obtained from pilot plant processing of non-hydrotreated gas oil with a density of 910 kg/m3 (56.8 lb/ft3). The pilot plant has a feed rate capacity of 500 – 1000 kg/hr (1102 – 2205 lb/hr).

Prefluidizing gas

Product gas

Feed Air

Regenerator stack gas

Figure 2.15: The NexCC process diagram

(after Ruottu et al., 2000).

2.9 Plasma Spouted Bed Reactor

A laboratory scale 2D plasma spouted bed reactor has recently been used to upgrade petroleum residue by researchers at the University of Paris. The high flow of hydrogen radicals is obtained by means of an inductively coupled argon-hydrogen plasma system as described by Leuenberger et al. (1993). The reactor is lined with a refractory material.

The radicals, introduced into the reactor laterally, tend to favor cracking reactions and therefore there is minimal coke formation. The spouted bed acts as a heat sink effectively quenching the plasma gas. The estimated quenching rate found was of the order of 5x105 °C/s. The plasma temperature may be as high as 3727°C (6740°F) while that of the spouted bed is around 1227°C (2240°F). The average residence time is given as 0.3 s.

The cracking yield obtained is on the order of 80% with a high selectivity of α-olefins and less than 0.1% coke formation with a feed of n-hexadecane. With a feed of gas-oil, the conversion is reported to be a 42%

Table 2.15: Representative results from the pilot reactor (after Hiltunen et al., 2002).

Run 1 Run 2 Reactor temperature, °C (°F) 599 (1110) 595 (1103) Regenerator temperature, °C (°F) 708 (1306) 741 (1366) Oil feed rate, kg/hr (lb/hr) 806 (1777) 1000 (2205) Reactor pressure, kPa (psia) 146 (21) 196 (28) Conversion, % 67.5 71.9 Yields, % dry gas 4.9 7.1 Ethylene 1.5 2.1 LPG 19.7 20.0 Propylene 6.7 7.1 C4 – olefins 8.8 8.6 gasoline (C5 - 221°C) 37.9 38.4 Coke 5.0 6.4

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reduction in the amount of heavy compounds (i.e. with a b.p. of greater than 300°C or 572°F) while the medium molecular weight components (150 < b.p. < 300°C or 302 < b.p.< 572°F) account for up to 40% based on the feed.

3 CATALYTIC CRACKING PROCESSES

3.1 Deep Catalytic Cracking

3.1.1 Introduction

Deep Catalytic Cracking (DCC) is a fluid catalytic process designed for selective cracking of hydrocarbon feedstocks into light olefins. Specifically, the DCC process was designed with heavier feedstocks in mind such as gas oils, vacuum gas oils, topped crudes, residual oils, tar sands, shale oil, and asphaltic fractions. The idea is to take advantage of the high activity catalysts that have been developed by increasing the operating temperature and decreasing the required residence required when using a standard FCC unit. The DCC design retains the steam stripping portion of the FCC design.

Riser reactor

Rams horn separator

Gas product and entrained solids

Stripper vessel

To downstream processing

Stripper riser (spent catalyst)

Gas-solids separator

Figure 3.1: The DCC process (after Letzsch et al., 1997).

The DCC process may be constructed as a new standalone unit, or may be incorporated into an existing FCC unit as an upgrade. Therefore, the two processes show a similar flow plan. The main differences include changes in the riser reactor and stripper configurations as well as changes in the operational temperature, partial pressures, and residence times. Research Institute of Petroleum Processing (RIPP) offers their own proprietary catalyst, but in principle any zeolite-type catalyst will function as far as olefin production is concerned.

The Deep Catalytic Process was a joint development between the RIPP and Sinopec. Stone & Webster Engineering Corp. (SWEC) has since been issued the sole license for offering this technology outside of the Peoples Republic of China. The first complex designed and engineered by SWEC was completed in Thailand in 1997 for Thailand Petrochemical Industries (TPI). In total, six such units are in operation with a combined feed rate of over 2x109 kg/yr (3x106 tons/yr) (shawgrp).

3.1.2 Process Description

Figure 3.1 illustrates the basic configuration of the DCC process. To begin, the hot catalytic particles are brought into contact with the hydrocarbon feed at the base of the riser reactor in a manner similar to an FCC riser reactor. No details of the feed mechanism are given. The dimensions of the riser reactor are chosen in order to achieve the desired residence time, typically on the order of 0.6 s. This residence time is based on the exit velocity from the riser. At the top of the riser reactor is located the solids separator as shown in Figure 3.2. Shown is a co-called “rams-horn” design that acts to divide the central stream of solids and vapour gas products into three separate streams. When the catalytic particles and vapour product stream enter the separator it

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first encounters the concave shaped central divider. The solid particles continue around the outer edge and exit in parallel conduits in a downward fashion. The gaseous stream exits from the circular outlets shown and are recombined before entering the stripper portion of the reactor system via the attached conduit. It is claimed that the portion of entrained solids remaining ranges from 0.1- to 2-wt % (Letzsch et al., 1997).

At the entrance to the stripper vessel the gaseous stream enters a second separation device, such as a cyclone, in the upper dilute phase of the stripper. The separated solids exit this cyclone and enter the dense phase found

in the stripper. Simultaneously the catalytic particles separated previously enter the stripper vessel via a separate conduit. Steam injections are used to pneumatically transport these solids. Steam is also used to carry out stripping of any remaining product vapours in the stripper vessel before the solids exit the stripper and are sent to the regenerator.

Figure 3.2: The DCC gas-solid

separator (after Letzsch et al., 1997).

Table 3.1: Comparison of typical yields for the DCC versus FCC

process (shawgrp). Product FCC DCC H2S 0.8 0.8 fuel gas 0.7 4.1 ethylene 1.5 5.4 propane 1.4 3.0 propylene 4.8 17.0 i-butane 3.2 4.3 n-butane 1.0 0.6 isobutylene 1.7 4.3 n,c,t-butylene 5.2 7.1 iso-amylene 2.7 4.2 naphtha 51.5 29.5 light cycle oil 15.1 10.4 heavy cycle oil 5.9 5.8 coke 4.5 6.6 Total 100 100 Feed: blend (50 % Arabian light and 50 % Brent vacuum gas oil by weight)

Finally, the product vapours exit the secondary separation device located from the top of the stripper vessel. From there they are sent to the quench system. No specifics are given as concerns the quench system used in the DCC process.

3.1.3 Results

A typical set of results for the DCC process versus a standard FCC process is given in Table 3.1. As may be seen from Table 3.1, there has been a marked increase in the production of fuel gas, ethylene, propylene, and isobutylene.

3.2 Total Raffinage Distribution

3.2.1 Introduction

Total Raffinage Distribution has developed a method for fluid catalytic cracking of hydrocarbon feedstocks based on studies that have led to a specially designed mixing chamber (Mirgain et al., 1998; Mirgain et al., 2000). The

emphasis of the design is on the optimization of the mixture quality as a function of the mixing chamber geometry.

3.2.2 Process Description

The flow plan for the Total FCC design is shown in Figure 3.3. In general, the reactor is operated in the downflow regime. Located at the top of the reaction chamber is the specially designed mixing chamber (shown in the square). A more detailed view of the mixing chamber is given in Figure 3.4. The regenerated catalyst particles enter the mixing chamber from above where they are brought into contact with the preheated hydrocarbon feed. The atomized hydrocarbon feed enters from several jets. The intimate mixture then descends downward through the reaction chamber before entering the stripper located at the base of the reactor. Just prior to the stripper chamber is

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Figure 3.3: The Total FCC flow plan

(after Fersing et al., 1999). Figure 3.4: The Total FCC mixing chamber

(after Fersing et al., 1999).

located the withdrawal line for the vapour products and the recovered vapours from the stripper. As the vapours pass through the horizontal conduit they are quenched with a hydrocarbon based liquid. Finally, the catalyst is transferred pneumatically using steam to the overhead regenerator.

In general, the mixing chamber geometry is such that there is a “completely agitated flow” that creates a large degree of back-mixing of the heated catalyst particles and the atomized hydrocarbon feedstock. It is this turbulence that leads to the intimate mixing. A detailed description of the mixing chamber geometry is given in US Patent 5, 997,726 (Fersing et al., 1999).

3.2.3 Results

The following example is discussed in the patent from Fersing et al. (1999). Table 3.2 gives the properties of the petroleum feedstock and the reaction conditions for a standard FCC reactor as compared to the Total FCC process

(shown in brackets). The results are presented in Table 3.3. Table 3.2: Petroleum feedstock properties and reaction conditions (after Fersing et al., 1999).

Density at 15°C (59°F) 0.925 50% distillation point, °C (°F) 470 (878°F) Viscosity, cSt at 100°C (212°F) 12.5 CCR, wt % 1.7 Nickel, wppm 0.1 Nitrogen, wppm 390 Vanadium, wppm 1 Catalyst Akzo

zeolite catalyst

Temperature, °C (°F) 520 – 545 (968 – 1013)

Number of injectors 8 Residence time, s 0.35 – 2

Table 3.3: Typical results for the Total-FCC process versus a standard FCC process (after Fersing et al., 1999).

Yield (wt %) Properties FCC Total-

FCC H2, CH4, C2H6 3.2 2.3 Paraffins (C3) 1.0 1.2 Olefins (C3) 3.3 5.9 Paraffins (C4) 1.9 2.6 Olefins (C4) 4.7 7.8 C5, b.p. < 160°C (320°F)) 31.8 34.8 Gasoline, b.p. 160 – 220°C (428°F) 11.7 11.5 Light cycle oil, b.p. 220 – 360°C (680°F) 19.1 17.2 Slurry, b.p. > 360°C 18.8 11.9 Coke 4.4 4.8

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3.3 Bartholic

3.3.1 Introduction

Several patents concerning short residence time cracking processes have been issued to Bartholic. The current state of development of all systems described herein is not certain, however it is known that at least one system accredited to Bartholic has been commercialized by UOP (refer to Section 3.4.2).

3.3.2 Solid-Liquid Hydrocarbon Contacting Process

This section describes a process developed by Bartholic (1996) for carrying out liquid-solid contacting for use in hydrocarbon catalytic cracking processes. The basis of this design is the designer’s claim that atomization of the liquid feed to droplets that typically range from 50 to 100 µm is not necessary. Rather, it is claimed that the distribution and surface area of the liquid feedstock exposed to the heated catalyst or inert solid is more important (Bartholic, 1996).

To this end, the Bar-Co process incorporates a series of nozzles that act to distribute the liquid hydrocarbon feed in what is described as a fan-shaped pattern of intermittent ligaments, or strings. This creates a large surface area through which the heated catalyst must pass thereby bringing the two streams into relatively intimate contact. Figure 3.5 (a) through (c) show the top, front, and side views, respectively, of the nozzle and of the spray pattern formed.

(a) (b) (c)

Figure 3.5: The (a) top, (b) front, and (c) side view of the Bar-Co nozzle (after Bartholic, 1996).

As may be seen for two different configurations in Figure 3.6 and Figure 3.7 this contacting scheme is for an upflow design. The former shows the design for a riser reactor where the solids enter from the bottom and rise through a curtain of liquid hydrocarbon droplets. The latter shows a slightly different configuration where the solids enter from the side at an angle and as they enter the upward stream of carrier gas they must again first pass through a curtain of hydrocarbon droplets. Finally, Figure 3.8 shows a top view of a riser reactor configuration involving a series of nozzles capable of introducing multiple feeds (i.e. a mixture). The figure shown corresponds to the riser configuration given in Figure 3.7.

Feed

Solids

Solids and vapors

Hot solids

Feed

Solids and vapors

Hot solids

Figure 3.6: Bar-Co solids-feed contactor (after Bartholic, 1996).

Figure 3.7: Alternative Bar-Co solids-feed contactor (after

Bartholic, 1996).

Figure 3.8: Top view of one configuration of the solids-feed mixer for the Bar-Co process

(after Bartholic, 1996).

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3.3.3 Horizontal Contact Scheme Cracking Process

Figure 3.9: Horizontal contact catalytic cracking

unit from Bartholic (1990).

Steam and/or hydrocarbons

Resid feed

Steam and/or hydrocarbons

Water

Gas Quench

Carbon and metals sorption

Figure 3.10: ART riser cracking process from

Bartholic (1981).

Figure 3.9 illustrates a horizontal contacting scheme for a fluid catalytic cracking process developed by Bartholic (1990; 1991). Heated fluidized catalyst is introduced into the reaction zone from an overhead standpipe vertically downwards to form a flat sheet with a velocity of approximately 6.1 m/s (20 ft/s). The “top hat” and the horizontal contactor form the reaction zone. The temperature of the catalyst is generally between 482 and 593°C (900 and 1100°F). The pressure within the system ranges from 69 to 345 kPa (10 to 50 psia). The catalyst to oil mass ratio of the system may be varied from 2 to 10. The zeolite catalyst used for this process is specifically designed to limit hydrogen transfer in order to provide better quality diesel fuel, less gas, and less secondary cracking.

Feedstocks include atmospheric and vacuum residual oils, gas oils, tar sand bitumen – feeds which typically have high metals, asphaltenes, and nitrogen contents. The feedstock is preheated to above 260°C (500°F). Atomization of the feed creates droplets ranging from 1 to 10 µm using gas with a nozzle pressure drop of between 3.4 to 207 kPa (0.5 to 30 psia) depending upon solids content of feed. The spray pattern created by the nozzle resembles that described in the previous section (i.e. a flat fan shape). The total contact time of this process falls between 0.1 and 0.2 s in the reaction zone plus time within the cyclone. Solids then enter the horizontal contactor at 15.2 to 30.5 m/s (50 to 100 ft/s). Following the contacting zone is a standard separation zone and product fractionation section. Solids that are recovered from the separation zone are returned to the bottom of the reactor vessel. Here, gas or steam is used to strip occluded gases from the surface of the solids. The flow rate of this gas may be varied in order to control the residence time of the catalytic solids entering the contactor. Excess solids are sent to a standard combustor for regeneration.

3.3.4 Riser Contact Scheme Cracking Process

The process described below was once referred to as the Asphalt Residuum Treatment process, or ART (see Figure 3.10). For this process, the typical feeds are resids with boiling points greater than 482°C (900°F) such as for vacuum resids or 260°C (500°F) in the case of atmospheric resids (Bartholic, 1982). Tar sand bitumen has also been treated using this process (Bartholic, 1989). For injection into the reactor system, the feed is atomized with steam and preheated to between 149 and 371°C (300 and 700°F). Hot inert solids, in this case specially prepared kaolin clay, enter the riser from a standpipe. The mixture temperature within the riser is on the order of (371 to 566°C (700 to 1050°F). The reactor pressure is maintained between 138 kPa (20 and 35 psia) (Bartholic, 1989). Compared to typical catalysts,

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the kaolin clay used has a relatively low surface area of 10-15 m2/g (4.9x104 to 7.3x104 ft2/lb), a corresponding low porosity, and a diameter of between 20 and 150 µm (Bartholic, 1982).

Velocity within the 6.1 m (20 ft) riser contactor is on the order of 12.2 m/s (40 ft/s) and the residence time is less than 0.5 s. The solids and product gases then exit the riser and enter the disengaging zone passing through several cyclones. Product gases exit from the top while the solids enter the stripper below and eventually are transported to a combustor for regeneration. The product gases, after leaving the cyclones, are quenched using cold hydrocarbon liquid. The bottoms are sent for further upgrading while the top product enters a condenser then an accumulator. The product liquid is claimed to be suitable for use as an FCC feedstock. No details

are provided on the exact means of introducing either the feedstock or the solids. It is stated that this process has been tested as a laboratory scale process using a circulating fluidized bed to simulate an FCC riser reactor (Bartholic, 1981) though this process strongly resembles a recently commercialized processes from UOP (refer to Figure 3.15 in Section 3.4.2). Properties of a typical feedstock are given in Table 3.4 along with a representative set results in Table 3.5.

Table 3.4: Feedstock properties (after Bartholic, 1982). Properties Feed 1 Feed 2 Gravity, °API 27.9 23 Ramsbottom Carbon, wt. % 0.35 2.5 Metals, wppm Ni 1 10 Cu 1 1 V 1 20 Distillation, °C (°F)

IBP 226 (438) 216 (420) 10 (%) 290 (554) 248 (478) 30 348 (659) 377 (711) 50 399 (750) 443 (829) 70 453 (847) 526 (979) 76 ( – ) 563 (1046) 90 533 (991) ( – ) 94 566

(1050) ( – )

Ashland Oil Company operated a 55000 bpd reactor in Catlettsburg, Kentucky beginning in 1983 (Milne, 1993). Table 3.6 shows a representative set of yields for this process. The feed is Utah tar sand bitumen with an API of 9.3° and 1.2-wt % mineral matter.

Table 3.5: Operating conditions and product characteristics (after Bartholic, 1982).

Operating conditions and product properties Feed 1 Feed 2 Riser contact temperature, °C (°F) 491 (915) 502 (935) Contact time, s 90.7 0.97 Solids temperature, °C (°F) 651 (1203) 641 (1185) Oil partial pressure, kPa (psia) 19.5 (2.83) 31.9 (4.62) Oil preheat temperature, °C (°F) 338 (641) 348 (659) Solids/oil, wt % 12.5 12.2 Mol. Ratio, N2/oil 3.7 2.2 Products, wt % Gas 7.9 7.6 Liquid 90.4 85.5 Coke 1.7 6.9 Liquid product Metal, ppm Ni ( – ) 1.5 Cu ( – ) 1.0 V ( – ) 1.0 Ramsbottom Carbon ( – ) 0.6 IBP 170 173 10, % 466 475 30 597 610 50 684 704 70 775 803 90 894 967 93 ( – ) 1033 EP 1028 ( – )

3.3.5 Downer Contact Scheme Cracking Process

Bartholic has also proposed a falling curtain design where the hydrocarbon is introduced into the reactor via a central

Table 3.6: Expected yields for ART process for a typical tar sand bitumen (after Bartholic, 1989).

LPG 2.7 C5 – 205°C 14.1 205°C – 345°C 10.8 345°C + 56.4 Coke 12.8

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Feedstock

Burner

Steam

Gas

Water

Annulus

Vaporization zone

FCC regeneratorFCC

reactor

Figure 3.11: Falling curtain type process from Bartholic (1984).

feed point and is surrounded by an annular ring of falling solids as shown in Figure 3.11. The inert solids, heated to between 593 and 871°C (1100 and 1600°F), form an annular sheet only a few inches thick. By comparison, the temperature of the hydrocarbon feed is between 371 and 454°C (700 and 850°F).

The feed is atomized with steam and sprayed with an angle of incidence as close to perpendicular as possible, but no less than 45°. The resulting mixture temperature reaches 482 to 566°C (900 to 1050°F). Finally, the reactor pressure is held between (69 – 103 kPa) 10 – 15 psia. The process is claimed to be an improvement upon the Houdresid design (Savage, 1951). While the design from Bartholic is said to maximize the production of the liquid components, the intention of the Houdresid design was to completely crack the feed to gases.

The product gases are withdrawn on the opposite side near the top of the vessel through uniformly spaced withdrawal lines in under 1 s. The hot gases are quenched using a cold hydrocarbon liquid while the coked solids form a dense phase fluidized bed at the reactor bottom, before being withdrawn and sent for regeneration. Steam entering from the vessel is used to strip any remaining product gas and maintain the required bed height needed to minimize the vapour residence time. The resulting synthetic crude oil is suitable for feed to a conventional FCC unit.

3.4 UOP

3.4.1 Introduction

Universal Oil Products (UOP), develops and licenses a large number of process technologies for the fluidized catalytic cracking (FCC) of hydrocarbons such as their Resid FCC (RFCC) process, Millisecond Catalytic Cracking (MSCC) process, and the PetroFCC process. As well, UOP has several patents relating to hydrocarbon feed injection systems (Optimix) and solid separators (vortex disengager stripper and vortex separation system).

All UOP catalytic cracking systems are based on conventional FCC designs and are intended either as revamp options or as new installations. In all cases the heat required for the process is supplied by regenerating the coked catalyst in a combustor. The RFCC process is targeted specifically for upgrading residual feeds with CCR values greater than 4-wt % using a two stage generator and a catalyst cooling stage. The PetroFCC process was designed to optimize the production of light olefins with propylene yields reaching 20- to 25-wt % depending upon the feedstock (UOP 2002 a). Finally the MSCC process, originally developed by BarCo (refer to Section 3.3) emphasizes short contact time catalytic cracking, something that is especially beneficial when cracking residual or heavy feeds. This process is described in greater detail in Section 3.4.2.

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3.4.2 Millisecond Catalytic Cracking

To date three Millisecond Catalytic Cracking (MSCC) commercial installations have been built with two more slated for completion in 2002. The first plant, built at Coastal’s Eagle Point Facility, has been in operation since 1994 (UOP 2002 a). The MSCC process is also marketed by UOP as a revamp option for existing FCC units to render them more suitable for heavier feedstocks and the benefits of shorter contact times.

The basic MSCC reactor configuration consists of an injection zone, a central dilute phase disengaging zone, the lower dense phase collection zone, and an upper inertial separation zone. Two basic reactor configurations for the MSCC process are illustrated Figure 3.12 (a) and (b), while Figure 3.12 (c) shows more detail of the injection zone (refer also to Section 3.4.3). Note the extended conduit with an opening to one side in Figure 3.12 (b), which forces the gas to make one full turn before entering the riser leading to the inertial separation zone.

To begin, catalytic solids at 593 to 760°C (1100 - 1400°F) enter from the upper conduit seen in Figure 3.12 (c) through the rectangular entrance thereby forming a dense “falling curtain” of solids (Palmas, 2000). The vaporized liquid hydrocarbon feedstock at 149 to 316°C (300 to 600°F) is then sprayed at 31 to 91 m/s (100 – 300 ft/s), in a fan-shaped pattern, through this curtain of solids. Total contact time is less than 1 s with a relatively high heat transfer rate (one estimate is 3.5x107 W/m3 (3x107 kCal/m3/h) from Schoenmakers, (1960)). The point of initial contact is situated near or just prior to the disengaging zone periphery. The cracked vapours are then withdrawn from the upper portion of the disengaging zone. It is possible to envision replacement of the catalytic solids with an inert heat carrier particle that could be used to crack heavy oil feedstocks. Figure 3.13 through Figure 3.15 show three previous processes to have incorporated a falling curtain contacting scheme, the latter dating back to 1960, while the former is from 1995. Figure 3.16 shows several more general arrangements. The QC process described in Section 2.3 could also be categorized as employing a falling-curtain contacting scheme.

(a) (b) (c) Figure 3.12: Part (a) and (b) show two different MSCC reactor configurations with (c) close of feed section

(after Palmas, 2000).

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Figure 3.13: FCC unit incorporating "falling-curtain" contacting scheme

(after Wegerer, 1995; Myers and Sattar, 2000).

Figure 3.14:Cracking process using falling curtain contact

scheme (after Schoenmakers, 1960).

Figure 3.15: FCC apparatus with falling curtain contact scheme (after

Radcliff and Cetinkaya, 2000; Lomas, 2001).

(a) (b) (c)

(d) (e) (f)

Figure 3.16:Additional falling-curtain contact schemes (after Schoenmakers, 1960).

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B B

A

A

A

B

Figure 3.17: Side and top view of two different inertial separation devices from UOP (after Niewiedzial, 1996).

In the UOP process, open volume is left above and below the initial disengaging section in order to allow the gas stream to rise unimpeded upwards. The solids removed at this stage enter the collection zone below where steam stripping can be utilized to remove any occluded gaseous products. In some design configurations the stripped gases may be immediately removed, or allowed to continue rising to the inertial separation zone. As well, hot regenerated particles may be sent directly to stripping zone to help with the stripping process by increasing the temperature within the stripping zone. Immediate separation of the solid and gaseous phases first occurs under the effects of gravity.

The gaseous products, and any entrained catalyst particles exit the disengaging zone and enter the inertial separation zone. A truncated cone, located just prior to the entrance of the inertial separator acts to increases the gas velocity. The inertial separator is designed with tangentially directed openings to create a centrifugal force on the particles as they exit (refer to Figure 3.17). Separated particles fall below to a fluidized bed before returning to the

collection zone via standpipes. The standpipes are typically submerged below the collection zone bed level to impede backflow and are spaced so as not to interfere with the incoming gas-solid mixture. Finally, the spent solids are sent to the regenerator after passing through a series of secondary cyclones, while the gaseous products continue for further conditioning and product recovery.

Table 3.7: Relative change in product properties for MSCC process vs. standard

FCC process (after UOP 2002 a). Product properties MSCC C2 (wt %) – 0.9 Gasoline yield (vol %) + 0.7 Liquid product (vol % + 1.7 Gasoline selectivity + 2.3

Schoenmakers (1960) emphasizes the importance of maintaining a continuous, dense vertical downward flowing curtain of solids, otherwise contacting is not as efficient. In other words, the solids flow should be fluidized or injected using a carrier gas. In Schoenmakers words, such a contacting scheme is “advantageous with astounding results” (Schoenmakers, 1960).

Table 3.8: Representative results from the process described by Schoenmakers (1960).

Molecular weight 366 Normal parrafins 35 Isoparaffins, naphthenes 63 Temperatures (°C) 750 Feed rate (kg/h) 40 Residence time (s) 0.4 Conversion (wt % feed) 80.2 Ethylene yield (wt % feed) 24.2 Methane yield (wt % feed) 8.2 Residence time (s) 0.4

A general set of results from such a contacting scheme as the one shown in Figure 3.14 is given in Table 3.7. Table 3.8 compares the results achievable with the MSCC process versus a standard FCC process.

3.4.3 UOP Optimix Feed Distribution System

The UOP Optimix feed distribution system is typically installed

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in an elevated position as shown in Figure 3.12 and Figure 3.13 orientated radially to the reactor vessel, and is suitable for either risers or downers. In contrast to Schoenmakers, the design of the acceleration zone below is such so as to provide a moderate catalyst density and velocity so that the solid particles are moving evenly through the contact region to allow for deeper penetration and mixing. In this fashion, it was reported that the formation of dry gas and coke were even less than in previous “falling-curtain” designs (UOP 2002 a). In addition, it is claimed that the UOP design uses less steam (e.g. 0.5- to 2-wt % of feed for a typical VGO).

As described above, the UOP feed distribution system is designed to create uniform thin flat fan spray pattern that is uniform over riser cross-sectional area. UOP has currently installed or licensed 64 of these systems for use in several dozen commercial applications (UOP 2002 c). Figure 3.18 shows an illustration of the UOP Optimix feed distribution system.

(a) (b) Figure 3.18: (a) top view and (b) side view of Optimix feed distribution system (Myers and Sattar, 2000).

3.4.4 UOP Vortex Separation Technology

The vortex separation system (VSS) was designed for internal riser reactors, while the vortex disengager stripper (VDS) is for external riser reactors. Both feature pre-strippers and are reported to be high efficiency (99.5%) (UOP 2002 a). It should be noted that UOP reports their efficiency as “percent hydrocarbon containment”. In other words, the measure how much of the gaseous product is removed from the solid and prevented from re-entering the reactor or remaining trapped on the solids. An illustration of the VSS configuration was given in Figure 3.15. The VDS arrangement is similar but with the solids arriving via an external riser (UOP 2002 d).

3.5 Exxon/Mobil/Esso

3.5.1 Mobil Mixer

Figure 3.19 shows a novel mixing tube developed by Mobil that is suitable for use on a fluid catalytic cracking unit riser. The tube comprises a venturi shaped draft tube. As shown, the catalyst recycled back to the base of the riser is split into two separate streams. The first stream is entrained into the centre of the mixing tube where it contacts the hydrocarbon feed. The second stream enters higher up through a series of slats located on the upper angled portion of the venturi draft tube. It is claimed that the venturi tube mixer helps minimize local overheating and reduces pressure surges (Chou and Lee, 1986).

3.5.2 Mobil annular falling curtain/plug valve

Mobil has also developed a falling-curtain type reactor for a short residence time catalytic cracking process as shown in Figure 3.20. The hot, regenerated solids enter the reactor from a solids surge vessel located directly overhead. As the solids fall they pass over a cone shaped plug valve that acts to create an annular curtain of solids. The plug valve is also connected to a vertical stem that is used to control the flow of solids (and hence the heat) in the upper portion of the reactor.

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Steam or nitrogen

Steam or nitrogen

Feed and steam

Catalyst/solids Feed injection

Stripping steam

Solids surge vessel

Fluidization gas

Feed injection or steam

Header (alternate)Plug valve seat

Plug valve

Bell separator (360°)

Vapor

Solids

Reactor products

Hollow plug valve stem

Figure 3.19: Mobil mixing zone

(after Chou and Lee, 1986). Figure 3.20: Mobil falling curtain with plug valve design

( after Raterman, 1994).

The hydrocarbon feed (atomized with 1- to 5-wt % steam) enters in one of two locations: either from the valve stem or the valve seat. It is also proposed that cooling fluid could be passed through the hollow stem in the valve when the unit is operated at higher temperatures. The hydrocarbon feed injected also partially coats the valve cooling it further. The typical reactor temperature is on the order of 760 to 820°C (1400 - 1500°F). The valve head is constructed of ceramic material. The vapour superficial velocity through the reactor is on the order of 6.1 to 61 m/s (20 to 200 ft/s) giving a total contact time of between 0.01 and 1.0 s.

The solids and gaseous products immediately enter the cylindrical bell-shaped separator located directly beneath the reaction zone. Within the stripper section there is a continuous upflow of stripping gas. The

Table 3.9: Estimated process yields for the Mobil falling curtain design (after Raterman, 1994).

Feed properties Base Case 1 Case 2 API ( - ) 24.0 ( - ) Con carbon ( - ) 0.42 ( - ) RI at 21°C (70°F) ( - ) 1.490 ( - ) Catalyst properties activity base Base 1.1 x base surface area, m2/g (ft2/lb) 185 (9.0x105) 185

(9.0x105)185

(9.0x105)

density, kg/m3 (lb/ft3) 1440 (90) 1440 (90) 1440 (90) Operating conditions Feed rate ( - ) Base ( - ) Feed temperature, °C (°F) ( - ) 262 (503) ( - ) Reactor temperature, °C (°F) ( - ) 521 (970) ( - ) Mix zone temperature, °C (°F) 536 (996) 533 (992) 532 (991) C/O ratio 9.7 14.2 14.0 Vapor residence time, s 3.04 0.12 0.11 Catalyst residence time, s 3.65 0.12 0.11 Yields, wt % dry gas 4.7 4.0 4.8 C2 1.4 1.2 1.5 C3’s 8.9 7.2 9.1 C4’s 18.0 14.6 18.4 gasoline 45.9 48.4 45.9 LCO 12.1 15.2 12.2 MCB 4.5 5.9 4.5 coke 4.5 3.5 3.6

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gaseous products then exit through the lateral outlet as shown.

Table 3.9 shows a set of estimated yields for the process in question. The base case represents the estimated yields for a conventional FCC riser. Case 1 then shows the estimated yields obtained using the process described above under similar operating conditions. Case 2 is similar to Case 1 but using a more active catalyst.

3.5.3 Mobil downer reactor

Another short residence fluidized bed catalytic cracker downer concept developed by Mobil is illustrated in Figure 3.21. Suitable feeds suggested include gas oils, and heavy petroleum fractions – in other words, typical FCC feeds.

In appearance, it is very similar to the Texaco design (see Section 3.6). The most striking difference is that in the Texaco design there is an initial separation of solids effected by gravity before the gaseous stream passes through a series of cyclones.

Flue gas

Air

Oil feed

Products

Stripper steamSpent

catalyst

Figure 3.21: Mobil downer reactor design

(after Gross, 1983).

As for the Mobil design, the entire gas-solid mixture is sent through the cyclones. Another similarity is the use of a circular perforated plate placed at the top of the downer reactor that acts to evenly distribute the falling catalyst particles. The hydrocarbon feedstock injection point is located level to, or just below, the catalyst introduction point. Most remaining components, such as the feed injection nozzles, stripper, and regenerator are conventional units. The downer itself is 37 m (120 ft) high and cylindrical with a diameter of 0.51 to 1.8 m (1.7 to 5.8 ft). In order to promote better mixing within the reactor, baffles may be installed at various points along its length. The superficial gas velocity is on the order of 4.6 to 11 m/s (15 to 35 ft/s) with a corresponding gas residence time of 0.4 to 2 s. Riser temperatures range between 530 and 650°C (980 and 1200°F).

Mobil claims that this particular process creates 20 to 30% less coke while producing 1% more gasoline than a conventional riser. Table 3.10 shows a comparative set of results between a conventional riser and the downer process described above.

Table 3.10: Representative yields for Mobil downer reactor (after Gross, 1983). Average Tmix=

524°C (975°F) Average

Tmix= 538°C (1000°F)

Average Tmix= 554°C (1030°F)

Oil residence time, s 5.29 5.18 5.25 5.31 5.21 5.40 Catalyst/oil, wt/wt 4.64 4.70 5.33 5.04 4.36 5.23 Calculated Tmix, °C (°F) 512

(953) 521

(970) 535

(995) 533

(993) 552

(1026) 554

(1029) Cracking yields, wt. % Conversion 62.10 61.08 65.20 64.61 61.34 65.62 C5/221°C (430°F) gasoline 42.83 42.22 43.15 44.15 43.30 44.41 C4’s 6.55 6.46 7.58 7.47 6.06 7.41 Dry gas 7.59 7.23 8.70 7.77 7.22 8.14 Coke 5.13 5.16 5.77 5.21 4.76 5.66

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3.5.4 Esso Short Contact Time System: Downer/Separator

Esso Research and Engineering Company has proposed a short contact time process for upgrading hydrocarbons that purposely accelerates the solids in order to effect a more rapid and complete separation from the product gas stream. As seen in Figure 3.22, the hydrocarbon feedstock, preheated to approximately 370°C (700°F), is sprayed into a downward fast flowing stream of hot particulate solids near the top of the reactor. Suitable feedstocks include heavy crudes, atmospheric vacuum bottoms, and pitch (Pappas, 1963). Also seen in Figure 3.22 is a series of slats and bounce bars that are used to effect an even distribution of the falling solids.

The solid particles are generally inert particles such as sand, fluid coke, or clay and may range in size from 100 to 3000 µm. The injection velocity of the hydrocarbon feedstock is such that the solid particles will be accelerated to at least 9 m/s (30 ft/s). Steam may also be injected if necessary to increase the solids velocity even higher. The temperature in the mixing zone is on the order of (650 to 870°C (1200 to 1600°F) while the mixture density is between 16 and 160 kg/m3 (1 and 10 lb/ft3).

While the mixing and reactor sections are both cylindrical and vertically aligned the separation section is curved. The radius of curvature is between 2 and 3 times that of the reactor diameter. As a result of the particle inertia, this curve acts to concentrate the solids along the outer radius. A baffle is also included that acts to concentrate the solid stream even further. The gaseous products escape laterally through a series of angled vanes that impede the flow of any entrained solids. The separation time is claimed to be under 0.25 s while the overall residence time is less than 1.0 s (Pappas, 1963). Immediately after the vanes the gaseous product stream is quenched. If necessary, another right angle turn in the direction of flow and another series of vanes can separate even more entrained solids. The solids are then collected in a hopper located further downstream. The gas stream is quenched one more time and then for downstream collection and processing. The solids are regenerated and then recycled back to the reactor. Finally, at the base of the downer a small reserve of solids is maintained by use of a weighted flapper valve in order to create a seal in order to prevent any

downward gas flow.

Steam

Oil feedOil feed

Hot solids

Reactor/ separator

Separator

Figure 3.22: Esso downer design

(after Pappas, 1963).

3.5.5 Esso Short Contact Time System: Riser

This design from Esso features two separate coking sections: the first is a dilute phase riser reactor and the second is a fluidized bed (see Figure 3.23). As to be expected, the gas residence time of the former is relatively short (1 to 2 s) compared to the latter. The riser is generally between 6 and 12 m (20 and 40 ft) in length.

The typical feeds include heavy residual oil with a Conradson Carbon number generally between 20 and 50. The feed is atomized with ½ mole of steam per barrel and injected at the base of the riser reactor. The mixture density is around 160 kg/m3 (10 lb/ft3), mass of coke to volume of gas, at riser bottom and 40 kg/m3 (2.5 lb/ft3) at the top. Coke particles serve as the inert heat carrier and are recycled back and forth from a fluidized bed combustor. The temperature within the primary riser is set around 900 – 1000°F. The solids and product gases exit the riser and immediately enter the secondary coking unit with the gases either passing up through the a dense phase bed as shown or through an initial separator (not shown).

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3.5.6 Exxon Short Contact Time System: Riser

A variation of the basic fluid coker process, Exxon has proposed a short residence time two-stage coker design as shown in Figure 3.24. In this case, the first stage is comprised of the cylindrical draft tube that is partially submerged within the fluidized bed, which in turn makes up the second stage. Other reactor concepts have included riser designs placed within the fluidized bed coker such as US Patent 2,885,272 (Kimberlin et al., 1959). In this design however, the draft tube was completely submerged within the bed of coke particles hence the residence time would be fairly similar to that of the standard fluid coker.

The fresh feed is introduced at the base of the draft tube along with steam and hot coke particles. Typical feedstocks can include heavy oils, residues, and bitumen. The temperature within the primary section is maintained between 540 and 590°C (1000 and 1100°F) and the residence time from 0.2 to 2 s. By way of comparison, the temperature in the secondary stage is between 480 and 520°C (900 and 975°F) while the residence time 20 to 40 s. The fluidization gas used within the secondary zone is steam. The reactor pressure is around 136 to 343 kPa (20 to 50 psia). The mixture velocity within the draft tube is between 11 and 18 m/s (35 and 60 ft/s).

At the riser outlet, the entrained solids are separated and fall into the secondary cracking zone. The gaseous products are sent downstream for processing with a portion of the heavier products recycled back to the secondary cracking zone. Table 3.11 summarizes a representative set of results for various residence times, temperatures, and feed types.

3.5.7 Exxon Short Contact Time System: Horizontal Bed

Exxon has also developed a fluid bed coker that utilizes a horizontal moving bed of fluidized heat transfer solids for the production of olefins. Typically the solids are coke particles that range in size from 400 to 800 µm, and may or may not be mixed with catalytic solids. The fluidization gas consists of gaseous hydrocarbons, hydrogen, hydrogen sulphide, and steam. In principle, the horizontal fluidized bed is as developed and described by Lurgi AG of Germany (see Section 2.7).

In the case of the Exxon process, the preferred feedstocks have boiling points that exceed 560°C (1040°F) (e.g. vacuum and atmospheric resides, bitumen, asphalt, tar sand oil, etc.). A portion of the

Gas Feed

Secondary coking zone

Figure 3.23: Esso short contact time system

(after Saxton, 1972).

Cold coke

Steam Hot coke

Hot coke

Fresh feed

Steam

Recycle oil

Scouring coke

Baffle

Figure 3.24: Exxon 2-stage design

(after Allan and Blaser, 1989).

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Table 3.11: Representative results for the Exxon 2-stage design (after Allan and Blaser, 1989). Feed 1 2 3 3 3 Temperature, °C (°F) 517 (962) 517 (963) 510 (950) 511 (951) 550 (1022) Pressure, kPa (psia) 186 (27) 200 (29) 173 (25) 171 (25) 193 (28) Steam injected, wt.% 11.2 11.4 14.1 30.5 29.3 Residence time, s 30.8 34.8 19.7 14.4 27.2 Ultimate yields C3 – gas 14.83 15.48 9.66 8.40 7.88 C4 3.18 3.00 2.28 2.08 1.81 C5 – 13.73 11.74 18.25 16.68 15.28 221°C (430°F) – 343°C (650°F) 8.68 7.34 14.40 12.95 13.54 343°C (650°F) – cut point 24.09 18.63 28.38 35.57 34.70 Coke 35.49 43.80 27.06 24.35 26.81

Feed 1 Vacuum residue 510°C (950°F) + recycle bottoms [API 0.4°, CCR 21.7 wt %]

Feed 2 Vacuum residue [API 6.7°, CCR 21.7 wt %] Feed 3 Heavy vacuum residue [API 7.8°, CCR 21.5 wt %]

heavy product may also be recycled back into the reactor as necessary. The heat carrier solids are heated to between 780 and 850°C (1440 and 1560°F). The vapor residence time is about 0.25 s, whereas for the solids it is from 10 to 30 s. Generally, the ratio of solids to feed within the reactor is maintained between 25:1 and 15:1.

The product is quenched using water or hydrocarbons streams, then introduced into a cyclone for separation of the solid and gaseous components. Typically, the product yield is 30- to 50-wt. % olefinic (5- to 15-wt. % methane, 10- to 30-wt. % ethylene, and 5- to 20-wt. % propylene). Table 3.12 shows a representative product yield obtained using a pilot plant that was fed South Louisiana vacuum residual oil at a feed rate 100 bpd. The

operating temperature is 396°C (745°F), and the vapour residence time is less than 1 s.

Table 3.12: Representative results from the Exxon horizontal bed design (after Serrand et al., 1997).

Feed rate, bpd 100 Temperature, °C (°F) 745 (1373) C3 - conversion Gas yields, wt. % feed 35 Gas Yields, wt. % feed Methane 7 –10 Ethylene 14 – 16 Propylene 9 – 12 Unsaturated C4’s 6 – 9 Liquid Yields, wt. % feed C5/220°C (428°F) 17.5 220°C (428°F)/340°C (644°F) 8.0 340°C (644°F) + 13.0 Gross coke, wt. % feed 18.7

3.6 Texaco Process

The Texaco process is another example of a fluid catalytic process adapted to a concurrent downflow regime. Figure 3.25 shows an illustration of the reactor and regenerator arrangement. In this case, the hot freshly regenerated catalyst in particulate form is undiluted (i.e. there is no carrier gas) and fed by gravity into the reaction zone. A catalyst distributor plate, as shown in Figure 3.26, is used in order to evenly distribute the catalyst prior to contact with the hydrocarbon feed. Texaco patented a similar concept at least as far back as 1959 as seen in US Patent 2,894,899 (Crawley, 1959).

The hydrocarbon feedstock is preheated to just below its vaporization temperature (e.g. 200 to 370°C or 400 to 700°F) prior to injection in the upper portion of the cylindrical reactor. There are three different injection ports aligned axially along the length of the reactor column equipped with several downward facing nozzles arranged radially around the outer wall of the reactor. Overall residence time of the system is adjusted by selecting the appropriate injection height. The reaction temperature is generally maintained between 510 and 570°C (950 and 1050°F), while the reactor pressure is on the order of pressure is between 130 and 310 kPa (19 to 45 psia).

The reaction products and spent catalyst immediately enter what has been termed a ballistic separator which utilizes the solid particles momentum to effect separation. In other words, the particles fall towards the

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bottom of the enlarged vessel while the gases escape laterally towards the cyclone train as illustrated in Figure 3.25 the entrance of which is level with the reactor exit. Located in the lower portion of the separator is a dense phase fluidized bed used for stripping. The stripper is maintained at a slightly higher temperature than the reactor itself. The stripping medium is steam. The flow rate of steam is such that there is essentially no net upward flow.

Flue gas

Oil feed

Product vapors

Stripping vapors

Regeneration gas

Figure 3.25: Texaco process diagram (after Niccum and Bunn Jr., 1985).

Figure 3.26: Solid feed distribution plate

(after Niccum and Bunn Jr., 1985).

Next, the spent catalyst enters the riser regeneration zone, is combined with air and burned, and finally discharged downwards from a similar ballistic separator into the solids surge vessel. The regenerator temperature is of the order of 680 to 790°C (1250 to 1450°F), with a residence time of 20 to 40 s, a density of 21 to 32 kg/m3 (1.3 to 2.0 lb/ft3), and a superficial gas velocity of 3 to 5 m/s (10 to 15 ft/s). The surge vessel also acts as the feeder from which reactor and stripping zone are fed.

No representative results concerning the product yields of such a process where located. The current status of this process is unknown. Niccum and Bunn Jr. make reference to a representative unit in US patent 4,514,285 (Niccum and Bunn Jr., 1985) that is 1.5 m (5 ft) in diameter and 12 m (40 ft) in length with a catalyst residence time of 1.7 s. The zeolite high activity catalyst entered the reactor at a rate of 372 kg/s (0.41 tons/s). The reactor inlet and outlet temperatures were 720°C (1325°F) and 530°C (985°F), respectively, while the reactor pressure was maintained at 275 kPa (40 psia). The products were observed to exit reactor at about 9 m/s (30 ft/s). The typical feed used in their study consisted of virgin gas oil with approximately 1.67-wt. % sulphur. The feed rate was the equivalent of 35000 bpd of fresh feed along with 1000 bpd of recycled product gas oil (Niccum and Bunn Jr., 1985).

4 OLEFIN PRODUCTION PYROLYSIS PROCESSES

4.1 Shock Wave Reactors for Olefin Production

4.1.1 Advanced Cracking Reactor (ACR)

The Advanced Cracking Reactor (ACR) was a joint development between Union Carbide, Kureha Chemical Industry Co., and Chiyoda Chemical and Construction Co dating to the mid 1970’s. The predecessor to this process is Kureha’s crude oil cracking process. The original development was towards an efficient process for producing olefins, especially ethylene. The economic incentive at that period was to move towards heavier feedstocks (naphtha, atmospheric gas oil, and vacuum gas oil) and away from the LPG’s (ethane, propane, and butane), which were anticipated to drastically increase in cost (Davis and Keister, 1976).

The process eventually selected is based upon high-temperature flame cracking originally developed by the two Japanese companies. The process involves mixing a stream of highly superheated steam with a stream of hydrocarbon liquids. Key to the high-temperature, low residence time, low hydrocarbon partial

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pressure process are the reaction system, quench and heat recovery system, and the cracked gas treating system. The technological development and assessment eventually led to the construction of a demonstration unit at Union Carbide’s Seadrift Texas plant (Baldwin and Kamm, 1983). The plant had a nominal ethylene production rate of 2.3x106 kg/yr (5x106 lbs/yr).

FuelEthyleneAcetylenePropyleneButadiene

Gas separators

Feedstock

Heater

Oxygen Steam

Burner

Reactor

FuelH2CH4

Quench oil

Heat recovery

Gasoline fractionator

Hydrocarbon oils

Water

Compressor 300 psig

CO/HS22

Acid-gas removal

Figure 4.1: The ACR process flow plan

(after Baldwin and Kamm, 1982).

Figure 4.2: Details of the ACR injection zone and

quenching zone (after Kamm et al., 1979).

However, the economic conditions were such that vacuum residue was eventually excluded as a feedstock because the yield of ethylene was insufficient for Union Carbide. However, the operating conditions of the ACR process are flexible enough that the range of propylene/ethylene ratio values is from 0.17 to 0.63 while the acetylene/ethylene ratio is adjustable from 0.04 to 0.13 according to Hu (1982).

4.1.2 ACR Process Description

The first stage in the ACR process is the burner-reactor (refer to Figure 4.1). The multiport burner is fed a gaseous fuel and pure oxygen mixture with steam injections used to moderate the temperature to 2000°C (3630°F). Next, the preheated hydrocarbon feedstock is injected in atomized form into the venturi shaped reaction chamber where “mixing, vaporization, and cracking occur simultaneously” (Baldwin and Kamm, 1983). The exact configuration of the nozzles and the spray pattern are given in Figure 4.2 as well as the upper portion of the quench chamber. At the point of feed injection the mean temperature is 1200°C (2190°F). From there, the mixture is accelerated to supersonic flow as it passes through the throat of the venturi. Next, the mixture passes through the stationary shock wave created just downstream of the throat where the hydrocarbon vapours undergo a second stage of cracking. The residence time in the reaction chamber is variable from between 15 and 30 ms. At this point, the products are at 900°C (1650°F) before being quenched to 335°C (635°F) in the heat recovery section. The reactor operating pressure is typically 446 kPa (65 psia).

The heat recovery section is also referred to as the Ozaki quench cooler (see Figure 4.3). As the cracked gas enters it is sprayed with quench oil. The quench oil is also used to form a continuous liquid film on the heat exchanger coils to eliminate tube side fouling. On the shell side, high-pressure steam at 11.8 MPa (1715 psia) is formed and used elsewhere in the plant. Makeup oil is drawn from the gasoline fractionator.

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The final step in the ACR process as designed by Union Carbide included a cracked gas treatment stage. Here, cracked gas is distilled and pyrolysis gasoline and fuel oil are recovered. Finally, after sulphur and CO2 removal, the cracked gas is dried and separated as product.

4.1.3 Typical ACR Results

Table 4.1 through Table 4.4 summarize the key results obtained using the ACR that are presented in the literature. Table 4.1 shows typical cracking yields for two standard feedstocks. Table 4.2 illustrates the flexibility of feedstock selection for the ACR process. The yields of cracked gases and naphtha are over 60 and 75%, respectively (Baldwin and Kamm, 1982). Table 4.3 shows the typical relationship between cracking severity and the resulting product yield. Finally, Table 4.4 shows the comparison for the olefin yields from naphtha and atmospheric gas oil feedstocks between the ACR process and a conventional olefin process.

4.1.4 University of Washington

The University of Washington has also patented a shockwave reactor system at the laboratory scale for the production of ethylene. The assumed product yield is upwards of 70% based upon an ethane feedstock. Figure 4.4 shows the process flow diagram. As may be seen, superheated steam is used as a carrier fluid. It is initially

Quenching oil nozzle

Lower portion of heat recovery section

High temperature gas introduction pipe

Quenching oil nozzle

Gas outlet

Heat recovery section

High pressure steam withdrawal port

Pre-cooling section

Distribution panel

Oil introduction port for forming oil film

Figure 4.3: The Ozaki quench cooler (after Hu, 1982).

Table 4.1: Typical cracking yields for the ACR process (after Hu, 1982).

Feed Arabian Light Crude

Pennsylvania Crude

Cracking temp, °C (°F) 960 (1760) 927 (1700) Residence time, s 0.015 0.030 Yield, wt % methane/hydrogen 11.23 11.3 acetylene 4.2 2.26 ethylene 31.8 34.0 propylene 6.1 12.0 butadiene 2.8 5.5 C5 – 220°C (430°F) 9.6 10.1 *220°C (430°F) + 24.1 12.0

Table 4.2: ACR feedstock flexibility, kg/kg feed (after Baldwin and Kamm, 1982). Feedstock Naphtha Gas oil

atmospheric Gas oil vacuum

Gas oil vacuum

Whole distillate

Products Ethylene 0.385 0.31 0.28 0.26 0.32 Propylene 0.16 0.11 0.11 0.11 0.12 Crude butadiene 0.10 0.08 0.08 0.07 0.08 H2/CH4/Fuel gas 0.14 0.12 0.12 0.12 0.12 Pyrolysis gasoline 0.17 0.20 0.11 0.11 0.20 BTX fraction (0.08) (0.10) (0.06) (0.06) (0.09) Pyrolysis fuel oil 0.06 0.18 0.31 0.33 0.18 Feedstock source Arabian

light Libyan Basrah Arabian

heavy Arabian

light API gravity 66.4 41.0 23.6 21.8 39.4 Boiling range, °C (°F) C5 – 149

(300) 149 (300) – 343 (650)

343 (650) – 566 (1050)

343 (650) – 566 (1050)

C5 – 566 (1050)

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Table 4.3: ACR product flexibility, kg/kg feed (after Baldwin and Kamm, 1982). Severity Low Medium High Products Ethylene 0.26 0.29 0.32 Propylene 0.11 0.10 0.06 Crude butadiene 0.08 0.08 0.05 H2/CH4/Fuel gas 0.10 0.12 0.16 Pyrolysis gasoline 0.15 0.14 0.14 BTX fraction (0.07) (0.07) (0.07) Pyrolysis fuel oil 0.30 0.28 0.28 Feed: Arabian light and Texas crude blend. 24.6 API gravity, boiling range 250 – 550°C (480 - 1025°F).

Table 4.4: ACR olefins yield for naphtha and atmospheric gas oil feedstocks (after Baldwin

and Kamm, 1982). C5 – 200°C (400°F) Naphtha 200°C - 320°C (400°F – 600°F)

atmospheric gas oil ACR Conventional ACR Conventional Feedstock 3314 4261 3283 4425 Products Ethylene 454 (1000) 454 (1000) 454 (1000) 454 (1000) Propylene 212 (467 333 (734) 171 (376) 308 (680) Crude butadiene (351) 228 (502) 131 (289) 229 (504) Pyrolysis gasoline (911) 571 (1258) 298 (658) 369 (813) Light fuel oil (69) 48 (106) 90 (198) 293 (646) Heavy fuel oil (118) 46 (101) 180 (397) 151 (333) Fuel gas (398) 254 (560) 166 (365) 204 (449) Feed: Typical Gulf coast domestic and low sulphur crude blend. Units:106 kg/yr (106 lbs/y).

Feedstock heater

Ethane feed

Carrier fluid heater

Nozzle section

Mixing Pyrolysis

Quench heat exchanger

Water feed

Ethylene product

Liquid

Figure 4.4: The University of Washington shock wave reactor system (after Hertzberg et al., 1994).

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heated to a temperature above that required for pyrolysis of the feedstock in question and pressurized to approximately 2.7 MPa (392 psia).

The carrier gas then enters the pyrolysis vessel comprised of the carrier gas inlet, a venturi section, a feed injection section, a mixing section, a pyrolysis section, and finally, the product outlet. The venturi section serves to functions: the first is to adiabatically accelerate the carrier gas to supersonic speeds (i.e. Mach 3) and increase the stagnation enthalpy while lowering the temperature so as not to adversely affect the mechanical components. The feed injection portion consists of a series of atomization nozzles for introducing the ethane feedstock in an essentially parallel manner to the direction of flow of the carrier gas.

The carrier gas and atomized feed mix in a turbulent manner due to the difference in velocity. Just upstream of the mixing section, the combined pressure of the atomized feedstock and carrier gas mixture induces a stationary continuous shock wave that causes a rapid increase in the mixture temperature to just above that required for pyrolysis. The products and carrier gas exit are then removed from the reactor at subsonic velocity. The duration of pyrolysis is a function of the position of the shock wave, which is controlled by adjusting the pressure of the feedstock, the carrier gas, and the back pressure induced at the outlet. Quenching of the product gas is carried out using a heat-exchanger or turbine. The overall residence time is on the order of 5 – 50 ms.

Table 4.5 shows the calculated energy inputs and outputs for the reactor system described above. As already stated, the expected ethylene yield is approximately 70% with the remaining by- product assumed to be hydrogen while products such as acetylene are ignored (Hertzberg et al., 1994).

Table 4.5: Energy input and output for University of Washington shock wave reactor system (after Hertzberg et al., 1994).

kJ/kg (kcal/lb) ethylene

Energy input: ethane heat (QE) 2400 (260) water heat (QW) 31820 (3447) Q = QE + QW total heat input (Q) 34220 (3707) ethane work (WE) 330 (36) pump work (WP) 70 (7.6) W = WE + WP total work input (W) 400 (43) Energy output: turbine work (WT) 6210 (673) process heating & cooling (QP) 23530 2549) Energy absorbed by reaction 4880 (529) Fractional conversion to ethylene 0.7 mass ratio steam/ethane 6.67 mass ratio steam/ethylene 10.20

4.2 Downer Catalytic Pyrolysis

4.2.1 Introduction

The Downer Catalytic Pyrolysis (DCP) process was designed for the production of light olefins from heavy feeds at Tsinghua University in China. The basic reactor design involves a hot downer reactor 4.5 m (14.8 ft) tall with an inner diameter of 13 mm (0.47 in). The overall height of the DCP system is 12 m (39.4 ft) with a maximum heating limit for the catalyst of 1200°C (2192°F). The downcomer design was chosen for its close approximation to plug flow. The standard hydrocarbon feed temperature is 350°C (662°F) fed at a rate of 3.5 kg/hr (7.7 lb/hr). The zeolite-based catalyst used in the process is to enhance the selectivity of the pyrolysis towards the production of ethylene and light olefins in general (Deng et al., 2002). Typically the operating temperature used is superior to 650°C (1200°F), which favors light olefin production and decreases the required residence time, as does a low hydrocarbon partial pressure. Figure 4.5 shows a simplified DCP process flow diagram.

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4.2.2 Process Description

To begin the catalyst is heated to 800°C (1470°F) before entering the mixer where it is sprayed with the preheated hydrocarbon feed. The pyrolysis and catalytic reactions occur as the components flow downwards towards the gas-solid separator. The solids then enter a stripper/regenerator where they will be regenerated prior to the next experiment. Meanwhile the gas enters a quenching vessel before being separated into gas and liquid products. No specific details are given concerning the nature of the mixer, gas-solids separator, or quench design (but it does appear similar to the configuration described by Mirgain et al., (1998) for catalytic cracking – refer to Section 3.2).

4.2.3 Results

Table 4.6 shows the basic properties for a typical feed to the DCP reactor. Table 4.7 then shows the test results obtained using the feed shown. Finally, Table 4.8 shows an analysis of the gasoline component from the DCP process where it may be seen that there is a relatively large portion of aromatics, which may create problems as regulations on gasoline are becoming more and more stringent.

Vent gas

Quench

Feed tankFeed

heater

Catalyst heaterCatalyst distributorFeed ejector

Gas solids mixerDowner

Gas solids separator

Gas product

Liquid productCoolers Stripper/

Regenerator

Figure 4.5: The Downer Catalytic Pyrolysis process

(after Deng et al., 2002).

4.3 Partial Combustion Cracking

4.3.1 Introduction

The partial combustion cracking (PCC) process is a development of Dow Chemicals originating in the mid-

Table 4.7: Experimental results (after Deng et al., 2002).

Reactor DCP Feed deasphalted oil Catalyst zeolyte Feeding temperature, °C (°F) 350 (662) Cracking temperature, °C (°F) 659 (1218) Operating pressure, MPa (psia) 0.005 (0.725) Catalyst/oil mass ratio 38.64 Residence time, s 0.75 Material balance Feed 100 Gas, wt. % 61.22 H2 0.45 CH4 9.16 C2H6 2.74 C2H4 16.06 C3H8 0.80 C3H6 24.57 C4H10 0.99 C4H8 10.91 Gasoline, wt. % 14.27 LCO, wt. % 12.74 Coke, wt. % 9.54 Loss, wt. % -2.23 Summary 100

Table 4.6: Feed properties (after Deng et al., 2002).

Component Carbon, wt % 12.85 Hydrogen, wt % 86.08 Sulphur, wt % 0.56 Nitrogen, kg/m3 (lb/ft3) 1525 (95.2) Density, kg/m3 (lb/ft3) 0.897 (0.060) Molecular weight 360

Table 4.8: Gasoline fraction composition from the DCP process

(after Deng et al., 2002). Components Aromatics, vol. % 73.5 Olefin, vol. % 23.2 Paraffin, vol. % 3.3

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1970s and continuing through the mid-1980s. The current status of development is not known since no reference to the DOW PCC process was found subsequent to this time frame. It is possible that development has been abandoned, is now operated under a new name, or that the technology has been sold to a third party. The last reported stage of development involved the construction of a 1000 bpd prototype unit in 1983 – 4 (Kirk, 1983).

4.3.2 Process Description

The DOW PCC utilizes combustion of a portion of the feedstock and recycled cracked feed to supply the heat of reaction required for the production of olefins by thermal cracking. Such feeds may include whole crude, fuel oil, and ‘heavy’ petroleum fractions (Kirk, 1983). The basic reactions involved in the process include partial oxidation in order to produce a reducing flame, a shift reaction where steam is injected to control the composition and temperature of the combustion gas, and finally a cracking reaction where the feed is introduced into the hot gas (Kirk, 1983). The process flow plan is shown in Figure 4.6. The main purpose of the shift reaction is to produce more hydrogen to aid in the conversion process and to convert the CO produced into CO2, which is more easily removed. Figure 4.7 shows an illustration of the reactor configuration along with the other principal components.

Shift zone Cracker

Quench and heat recovery

C + recovery 2 Compression Primary fractionation Acid gas recovery

Product recovery

Acetylene conversion

Olefins Fuel gas (Methane + syn-gas)

Acetylene recovery

SulphurHS conversion 2

Burner

CO 2

Steam Metals, Pitch

Steam Steam Feedstock

Reactor

Oxygen Recycled fuel oil and gas

V

V

V

V

V

Product gases

Feedstock

Oxidizing gas

Fuel

Steam

V

VSuperheated steam

Figure 4.6: The DOW PCC flow plan (after Kirk, 1983). Figure 4.7: DOW PCC reactor configuration

(after Read et al., 1981).

The reactions are carried out at pressures ranging from 35 to 448 kPa (5.1 to 65 psia) and temperatures of 1600 – 2000°C (2912 – 3632°F), 1300 – 1600°C (2372 – 2912°F), and 700 – 1000°C (1292 – 1832°F) for the combustion, shift, and cracking reactions, respectively. The ratio of feed to fuel typically ranges from 1.5:1 to 6:1 in order to obtain a residence time on the order of 0.05 to 0.5 s (Read et al., 1981). The product gases are sent to what is described as a “unique” heat exchanger for rapid quenching – although no details are given. The remaining treatment and recovery stages make use of conventional process technology. Mention is also made of a specially designed refractory lining for the adiabatic reactor, PCC burner, and the feed injection nozzles (Kirk, 1983).

4.3.3 Results

Typical ethylene yields using the PCC process are typically around 24 to 35% as shown in Table 4.9. Table 4.10 shows the overall cracking yields for three different feeds.

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4.4 Arthur G. McKee & Company Table 4.9: Ethylene yields for the PCC process as compared to standard tubular technology (after Kirk,

1983). Feedstock PCC Tubular

technology IBP – 177°C (350°F) 35 % 28 % 177°C (350°F) – 343°C (650°F) 32 % 23 % 30 API crude oil 27 % - Residual oil 25 % -

United States Patent 4,172,857 (Pavilon, 1979) describes a thermal cracking process for the production of ethylene by the pyrolysis of ethane, gas oil, naphtha, residual oils, or tar sands developed by Arthur G. McKee & Company (now part of the Aker Kvaerner Group). The pyrolysis unit is a riser reactor where hot agglomerated ash particles are used for the heat source. Incorporated into the process is a fluidized bed burner used to recycle the hot agglomerated ash by burning off the accumulated char and generate hot flue gases that are used to create superheated steam and preheat the gaseous or liquid feedstocks.

In general, the process flow plan for this process exhibits many similarities with other designs. Hot char particles are transported through the cylindrical riser using superheated steam. Next, preheated gaseous or liquid feed is sprayed upwards into the column along the vertical axis of the pyrolysis column. Temperatures in the riser reactor are maintained between 700 and 1150°C (1292 and 2102°F) and the overall residence time is under 1 s (Pavilon, 1979). A preliminary separation of the gas stream from the particulate stream in a rough cut cyclone located at the exit of the riser. Here, these components are quenched using water. These jets are

described as being directed upwards towards the centre of the vessel in order to avoid the hot ash particles that would generally be found towards the outer wall. If this were not the case, then a certain amount of energy would be wasted in quenching the hot solids that would then require even more energy to be reheated. The gaseous stream then passes through another train of cyclones before entering the product recovery section.

Table 4.10: Typical cracking yields for the DOW PCC process, kg/100 kg feed (after Kirk, 1983).

Products Crude oil VGO 343°C (650°F)

residual oil Methane 14.0 12.5 12.4 Acetylene 2.3 2.7 2.5 Ethylene 25.2 24.5 24.4 Ethane 1.9 1.2 1.6 MA/PP 1.0 0.7 0.7 Propylene 8.3 5.5 6.9 C4’s 6.6 2.9 5.1 C5’s 3.0 1.2 1.6 C6 – C7 20.2 17.8 14.3 C8 + fuel oil 15.7 27.3 28.5 H2S 0.4 1.7 0.3 Balance: mixture of CO/CO2/H2 depending on reaction severity.

Figure 4.8: Simplified flow diagram for the “McKee” process

(after Pavilon, 1979).

The coke covered ash particles are then sent to a fluidized bed stripping unit where steam removes any occluded gaseous products trapped in the solid particles. The ash particles, still covered in coke, are then sent to the burner to complete the loop. A simplified process diagram is given in Figure 4.8 showing the fluidized bed agglomeration burner (to the left), the rough-cut cyclone and stripper (centre), and the riser reactor. A more complete description of the process flow plan is given by Pavilon (1989).

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4.5 Ube Process

Ube Industries started developing the Ube process in the 1960’s eventually constructing a 4535 kg/d (5 tons/d) pilot plant in 1969 followed by demonstration scale plant in 1979 operating at 230000 kg/day (250 tons/d). The process, as described by Hu (1982), consists of a fluidized bed of hot mullite particles that serve as the heat source. The fluidization gas consists of steam and oxygen, the latter of which is used to promote partial combustion of the hydrocarbon feed that provides the required heat for the process.

Methane

CO2

H2

Demethanizer

CO and C H separation

2

2 2Desulphurizer

Ethylene Propylene C fractions3

C fractions4

C /C /Cfractions separator

2 3 4

Heavy oil

Water

Cr. Gasoline

Gas oil

Distillation column

Flue gas

Regenerator

Air

Fuel oil

Quencher

Steam

Oxygen

Crude oil

Fluidizing gas - steam

Gas jet - steam

Water or light hydrocarbons

Figure 4.9: Ube process flow diagram (after Hu, 1982).

Feed to the bed is described as crude oil that has been preheated to 400°C (750°F) and blown into the bed. It is not known if the feed jet is horizontal or vertical, but Milne (1993) does describe the process as a spouted bed. The temperature within the bed is maintained between 832 and 879°C (1530 and 1615°F). Velocity of the gas jet is between 25 and 35 m/s (82 and 115 ft/s), which corresponds to a residence time of between 0.2 and 0.3 s. The fluidization velocity is approximately 6 to 7 m/s (20 to 23 ft/s). The coked mullite particles are removed periodically from the fluidized bed and sent for regeneration while the product gases are quenched by direct contact with water or light hydrocarbons. Figure 4.9 shows a general process flow diagram for the Ube process while Table 4.11 shows a set of typical process yields.

Table 4.11: Typical process yields for the Ube process (after Hu, 1982).

Feed crude oil Temperature, °C (°F) 840 (1544) Residence time, s 0.2 – 0.3 Yield, wt % ethylene 28.1 propylene 11.3 C4 fraction 8.4 gasoline 12.2 heavy oil 4.4 gas 66.6

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4.6 COSMOS Process

Development of the COSMOS (cracking oil by steam and molten salts) began in the early 1970’s. The molten salts (e.g. LI2CO3, Na2CO3, and K2CO3 mixtures), also referred to as the melt) are used as the heat carrier. These salts catalyze water-gas reactions and help prevent plugging by impeding coke buildup on the reactor wall – the so-called wet wall effect.

The COSMOS process makes use of a conventional externally heated tubular reactor. High velocity steam, heated to 750°C (1382°F), is injected into the reactor where it mixes with the preheated feedstock. Typical feedstocks may include crude oil or heavy oil, or anything that may be paraffin rich but low in sulphur due to the tube metallurgical limits. For injection the feed is atomized using steam.

Table 4.12: Feed properties (after Yamaguchi et al., 1979). Feed Minas

crude Minas crude

bottoms Taching

crude Specific gravity at 45°C (113°F)

0.838 0.88 .842

S, wt. % 0.12 .31 .17 Distillation, °C (°F) IBP 103 (217) 295 (563) 70 (158) 5 % 310 (590) 160 (320) 10 % 171 (340) 317 (603) 203 (397) 30 % 302 (576) 325 (617) 335 (635) 50 % 35 % at

350 (662) 40 % at 370

(698) 37 % at

350 (662) C, wt. % 84.46 86.81 85.56 H, wt. % 13.43 12.48 13.45 CCR, wt. % 2.5 3.4 2.7

Table 4.13: COSMOS process yields (after Hu, 1982; Yamaguchi et al., 1979)

Feed Minas crude Minas crude bottoms

Taching crude

Temperature, °C (°F) 800 (1470) 800 (1470) 800 (1470) Pressure, kPa (psia) 101 kPa (14.7) 103 (14.9) 102 (14.8) Feed rate, kg/hr (lb/hr) 3 (6.6) 3 (6.6) 3 (6.6) Steam/oil ratio, wt./wt. 1.33 1.33 1.33 Residence time, s 0.55 0.54 0.55 Melt/oil ratio, wt./wt. 2.01 ( - ) 2.09 Yield, wt. % H2 + CH4 15.71 14.81 14.56 C2’s 28.28 26.7 27.7 C3’s 11.88 11.57 11.35 C4’s 6.11 6.09 6.65 C5 + 9.31 7.92 8.09 CO2 11.75 17.84 15.30 Fuel oil 25.86 29.02 28.39 Others balance balance balance

Product gases, melt, and steam exit reactor the and are quenched then separated using a cyclone with the melt being partially recycled back to the reactor. Table 4.12 shows the properties for a typical feed and Table 4.13 a representative set of results obtained using a 5 kg-oil/hr (11 lb/hr) bench scale unit, respectively. Figure 4.10 shows a process flow diagram of the COSMOS process. The current state of this process is not known.

Melt separator

Melt drumMelt pump

Cracker

Quencher

Dilution steam

Crude oil atomizing steam

Recycle melt

Cracked gas

Figure 4.10: COSMOS process diagram (after Hu, 1982).

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4.7 Paccal Process

4.7.1 Introduction

The Paccal process is a development of The Petroleum & Chemical Corp used for the production of olefins, particularly ethylene, from crude residuum feedstock. The first process unit was built in 1955 at Silverwater in New South Wales, Australia. Production at this facility continued up to 1974 when the low relatively low cost of natural gas made the Paccal process economically unviable and the plant was forced to shut down.

4.7.2 Process Description

The Paccal process involves the thermal cracking of atomized crude that is sprayed downwards onto a stacked bed of hot refractory checker bricks in an essentially vertical refractory lined reactor. The reactors, six in total, were operated in pairs in a cyclical reaction/regeneration pattern approximately four minutes in duration. First, atomized feed and atomization steam were introduced into the reactor and cracking proceeded for 1 ½ minutes. Next, purging steam was introduced into the reactor along with fuel (oil or gas) and regeneration air, and finally purging steam. The reactors themselves were 4 m (13 ft) in diameter, 15 m (50 ft) high with 6 m (20 ft) of checker bricks placed in the lower portion. These bricks were maintained at 1000°C (1832°F). Changing the brick bed height controlled the vapour residence time. The pressure in bed was approximately 203 kPa (30 psia).

4.7.3 Results

Upon exiting the reactor the product gases were quenched using water and sent for further processing. Yearly production rates reached 30000 metric tons/y of ethylene, 17000 metric tons/y raw BTX, 16000 metric tons/ aromatic distillates, 17000 metric tons/y petroleum pitch and a variety of tars. Suitable feeds included tar sands and shale but higher C/H ratios tended to favor an increased pitch production while ethylene production would decrease. The gases produced were essentially 30-wt. % ethylene and 9-wt. % propylene (Cusack and Mashford, 1974). Table 4.14 shows a representative set of results for the processing of several different crude residuum feeds.

Table 4.14: Yields for the Paccal process for several different feeds, wt % of feedstock (after Cusack and Mashford, 1974).

Feedstock Australia (Gippsland)

Indonesia (Minas)

Middle East (Iranian)

FCC cycle oils

Gravity (cp) 0.88 0.89 0.975 0.92 C/H 6.4 6.3 7.7 7.3 Paccal yields hydrogen 1 1 1 1 dry fuel gas 22 21 16 18 ethylene 21 20 15 16 propylene 7 7 5 5 1,2 butadiene 2 2 1 1 butenes 2 2 1 1 C5/C6 unsats 2 2 2 2 benzene 7 7 5 6 toluene 3 3 2 3 xylene 1 1 1 1 arom. distillates 11 13 15 18 pitch 13 13 22 18 carbon 8 8 14 10

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5 BIOMASS FAST PYROLYSIS

5.1 Biomass Technology Group (BTG)

5.1.1 Introduction

Originally part of the Chemical Engineering Department at the University of Twente, the Biomass Technology Group (BTG) b.v. is now an independent private firm. BTG activities centre upon biomass energy technologies including pyrolysis, carbonization, gasification, and combustion (Bridgwater and Peacocke, 2000).

The principal method developed by BTG for the pyrolysis of biomass centres upon their rotating cone technology. Development began in 1989 as a PhD thesis project at Twente University (Bridgwater and Peacocke, 2000). This work resulted in the development of a 10 kg/hr (22 lb/hr) prototype reactor. Subsequent projects included another PhD thesis to develop a fully heat integrated pilot plant in 1994 and a means of performing catalytic pyrolysis in 1996. Now, scale-up work for pilot plant design and commercialization is underway between BTG and Royal Schelde and Kara. This has resulted in a 50 kg/hr (110 lb/hr) test unit being constructed at the Shenyang Agricultural University in China (Bridgwater and Peacocke, 2000). Scale-up beyond a reactor 2 m in diameter is carried out by stacking several cones one on top of the other and aligning along a single axis. In this manner, pilot plants may be built that have feed rate capacities of 1800 to 90000 kg/hr (2 to 100 tons/hr). A simplified flow diagram of a 200 kg/hr (440 lb/hr) unit is shown in Figure 5.1. Clients may have their feeds tested at the BTG facility.

5.1.2 Process Description

As may be seen in Figure 5.2 the BTG reactor has a conical shape; the wide end towards the top that then tapers downwards towards the smaller end at the bottom. Attached to the upper portion of the housing, at the tapered end, are feeding points for the heat carrier (e.g. sand) and the non-gaseous biomass feed (e.g. sawdust). The biomass feed tube is cooled by water. A third port serves as an outlet for the pyrolysis vapours that are produced. Figure 5.3 shows a 3D illustration of the reactor configuration. The entire conical reactor is housed in a furnace.

For the configuration shown, the lower portion of the conical body would rotate round its longitudinal axis while the upper portion would be fixed in place. The rotation of the lower portion exerts a centrifugal force on the non-gaseous biomass feed and the solid heat

carrier found in the bottom volume. These components will migrate towards the outer edge of the bowl and up the inner wall within the annular volume. Rapid and intimate mixing and heat transfer occurs between the heat carrier particles and the non-gaseous biomass occurs within the bottom volume of the reactor. A high rate of heat transfer also occurs at the heated cone surface as there is relatively close contact between the particles and the rotating cone. Any non-gaseous components remaining after pyrolysis are ejected at the top of the conical reactor into the dead volume as a result of the centrifugal forces created. Tracer testing indicates that there is negligible vapour transfer to the dead zone volume. The company GEM has developed a similar system. However, rather than having the solids migrate up a conical insert, centrifugal force is exerted on the particles by means of a fan located at the top of the reactor vessel and the particles of biomass migrate downwards instead (Maton, 1999).

Ash

Saw dust

R1

R2

Flue gas

Air

Air

Gas

C1

H1

Oil

Water

C2

Figure 5.1: The process flow plan for the BTG rotating cone

pyrolysis system (after BTG, 2002).

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Pyrolysis vapors

Char and sand

Reactor volume:annular volumeBottom volume

Blocked inner cone volume

Sawdust

Heat carrier (sand)

Dead volume

Rotating cone

Pyrolysis vapors

Sawdust Preheated sand

Figure 5.2: Schematic of the BTG rotating cone pyrolysis system (after Wagenaar et al., 1994).

Figure 5.3: 3D representation of the BTG cone reactor (after Wagenaar et al., 1994).

There are two main advantages of the rotating cone system. The first is the elimination of the requirement for a carrier gas, which helps reduce the cost and increase the product gas concentration. The second is the reduction in the required reactor volume, which reduces the gas-phase residence time (Wagenaar et al., 1994). Difficulties may arise however, if inaccurate estimations are made of the feedstock residence time since this would affect the product distribution. Other difficulties include the operation of three complex systems including the rotating cone, the solids riser to the burner, and the burner itself (Bridgwater, 2001).

The product collection system for the study conducted by Wagenaar et al. (1994) consisted of a cyclone followed by six washing bottles, and a filter. The washing bottles were placed in an ice bath in order to condense out the liquid tar (or bio-oil) that was formed using the sawdust feed. The filter was designed to collect any product vapour.

As concerns the solids circulation loop, two different configurations have been proposed. In the first, as the solid particles (heat carrier and accumulated char) exit the upper portion of the conical reactor they enter the dead volume. This portion, within the enclosed furnace, is fluidized to maintain easy flowing solids. The rotating cone is then equipped with a scoop that allows the particles to reenter the reactor (refer to Figure 5.4). In the second configuration, deflector plates redirect the solid flow back to the centre of the rotating cone. The solids then enter the inner volume via an opening between the feed port (refer to Figure 5.5). In either case, escaping solids are transferred to a regenerator where the char is combusted in order to reheat the sand particles, which are then fed back into the reactor (refer to Figure 5.1).

Figure 5.4: Rotating cone with scoop

(after Wagenaar et al., 2001). Figure 5.5: Rotating cone with deflector plates

(after Wagenaar et al., 2001).

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5.1.3 Results

In the study by Wagenaar et al. (1994) experiments were conducted for a range of temperatures from 550 to 700°C. The reactor volume was varied from 0.003 to 0.2 m3 and the cone rotational frequency between 6 and 16 Hz. Typically, each experiment lasted 8 min and between 0.5 and 2.0 kg (1.1 and 4.4 lb) of sawdust and 15 kg (33 lb) of sand was fed into the reactor. The maximum tar yield achieved was 70% (on dry wood basis). The results show an increase in the biomass conversion for increasing temperatures and a decrease in the rotational frequency. The stated residence time of the system is on the order of 1 s with the reactor vapour residence time being 0.3 s. The solid residence time is about 0.5 s. A typical set of results for an analysis using pine sawdust is given in Table 5.1. It is also estimated that less than 10-wt% of the product vapour is lost due to secondary cracking (BTG).

Table 5.1. Typical results for the pyrolysis of pine sawdust using the BTG rotating cone (after SciTecLibrary.com, 2000).

Biomass feed rate, kg/hr (lb/hr) 10 (22) Rotary speed, rpm 240 Liquid yield, wt % 50 – 60 % Gas yield, wt % up to 15% Gas products CO, CO3, CH4, and H2Solid yield, wt % 10 (max)

5.2 Dynamotive Technologies Corp. and Resources Transforms International Ltd.

5.2.1 Introduction

Dynamotive Technologies Corporation (hereinafter referred to as Dynamotive) was incorporated in 1991 and is located at BC Research in Vancouver, British Columbia. Their technologies centre upon the development of fast biomass pyrolysis systems and the derivation of value added products obtained from the bio-oil.

Beginning around 1997, Dynamotive entered into a partnership with Resources Transforms International Ltd. (or RTI) based in Waterloo, Ontario. The result of this collaboration has been the development of what is referred to as the BioTherm™ process (Bridgwater and Peacocke, 2000). The BioTherm™ process is a fluidized bed pyrolysis system. The first unit constructed in 1997 achieved a production capacity of 80 kg/hr (176 lb/hr) utilizing a mixed sawdust feedstock (Bridgwater and Peacocke, 2000). Currently, the focus of Dynamotive is set on commercializing their biomass refinery technology through partnerships with Lockheed Martin/INEEL, Usina Santa Helena SA, and Stone and Webster Engineering Corp.

5.2.2 Process Configuration and Development

The original process technology developed by RTI was the result of studies conducted at The University of Waterloo. This process was dubbed the Waterloo Fast Pyrolysis Process (WFPP). At the heart of the WFBB system was a shallow fluidized bed operating in the bubbling regime. Typically, a relatively high gas flow rate is required in order to maintain the short residence time sought (e.g. ~ 0.5 to 1 s).

Although shallow fluid beds may be readily scaled-up the limit is quickly reached where height-to-diameter ratio will lead to temperature and concentration gradients. As a result of the higher gas flow rate special attention is also required in the design of the downstream processing equipment. A decrease in the process thermal efficiency also occurs since the recirculating gas must be cooled as it passes the condensers and then reheated before reentering the bed. It is estimated for example, that the thermal efficiency of the WFPP system is between 60 and 70% (Scott et al., 1999). However, the liquid product produced is of high quality with a minimum of char (solids) formation. The char formed does not tend to accumulate in the bed; instead it is readily elutriated from the bed.

An industrial example of the WFBB system exists in the form of the Union Fenosa plant in Spain. Typically, the shallow fluidized bed of sand is operated at 500°C (932°F), 151 kPa (22 psia), with a 0.5 – 1 s residence time (Piskorz et al., 1998). The production capacity of the plant is 200 kg/hr (440 lb/hr). In this situation the limitation in the thermal efficiency manifests itself since the heat delivery is coupled to thermal capacity of the recirculated gas. The temperature of the returning gas is kept below 900°C (1652°F) to impede coke formation. Therefore, any increase in the thermal capacity must come from an increase in the gas flow rate with the subsequent increase in loading for all downstream equipment (Piskorz et al., 1998).

However, the requirement for such relatively high temperatures and short residence time has been questioned in recent studies by the RTI group (Scott et al., 1999; Piskorz et al., 1998). In fact, they claim equivalent

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liquid quality and quantity are obtainable with the use of a deep fluidized bed operating at lower temperatures and higher residence times (upwards of 10 s). Table 5.2 in Section 5.2.4 shows a simple comparison between the WFBB process (shallow bed) and the BioTherm™ process (deep bed) as well as the ENSYN/RTP process (refer to Section 5.3). In all cases, the production of char using the BioTherm process is considerably less at comparable temperatures.

The use of the lower operating temperature allows the additional advantage of being able to employ standard heat exchanger tubes within the fluidized bed, although any practical heating method may be used. As well, standard deep fluidized beds offer simple construction and operation, good temperature control, efficient heat transfer, and easy scaling since it is a fairly well understood technology (Bridgwater, 2001). With such technology, good char separation is necessary and is usually accomplished with one or more cyclones. This is because one concern with such technology is contamination of liquid product by the solid residue. Typically the bed particles are less than 2 or 3 mm (0.08 or 0.12 in) in size in order to facilitate biomass heating.

5.2.3 Process Description

To begin, the prepared feedstock for the BioTherm process, usually wood sawdust, is conveyed by auger from the storage hopper to the metering feedbox from which a constant speed injector screw feeds it into the reactor (refer to Figure 5.6). Heat is supplied to the reactor via a natural gas burner and external heat exchanger jacket. Once inside the reactor the biomass feed is quickly heated to the typical operating temperature of 450°C (842°F) (BioTherm, 1999). The standard bed material used for the process is sand, while the standard products are of course char, gas,

vapours, and aerosols, which then exit the top of the reactor before entering a series of cyclones. The entrained solid particles are then simply collected in reservoirs located beneath the cyclones.

Following the cyclones is the recovery direct quench system where the gaseous products are cooled to below 50°C (122°F) using an immiscible liquid (BioTherm, 1999). The condensed bio-oil is collected and the coolant separated and recycled back into the process. Finally, an electrostatic precipitator is used to remove any remaining aerosols from the gas. The inert gas, now cleaned, is recycled back to be used for fluidizing the system. Generally, the non-condensable product gas is sent to a flare, however, design modifications are under way to send this gas to the process burner instead.

Feedstock hopper

Feed box

Feeder

Biotherm reactor

Burner

Natural gas

Combustion air

Cyclones

Char pots

Cooling water

Quench liquid pump

Quench system

Electrostatic precipitator

Bio oil product

Product tank

To flare

Feedstock conveyor

Figure 5.6: The BioTherm process flow plan

(after BioTherm, 1999).

5.2.4 Results

Table 5.2 shows some general results for the BioTherm process. In general the results are reported to be very consistent with liquid yields between 70 – 75-wt % from wood on a dry basis. The results have also been shown to be very consistent during the scaling up from laboratory scale to the large pilot plant.

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5.2.5 Fast Pyrolysis of Waste Plastics

Similar to the ICFB process (see Section 2.1), studies are being conducted using the Waterloo Fast Pyrolysis Process

(WFBB) (i.e. the Dynamotive process) in order to determine its suitability for fast pyrolysis of waste plastics. Studies have been conducted using a bench-scale (feed rate of up to 0.10 kg/hr or 0.22 lb/hr) and a pilot plant (feed rate of up to 5 kg/h or 11 lb/hr) using feeds of poly(vinyl chloride), polystyrene, and polyethylene. The relatively low melting point of the polymer feeds compared to that of the biomass required modifications to the feeding mechanism. Whereas the biomass was fed by means of a screw feeder it was necessary to introduce the polymer feeds from the top of the reactor.

Table 5.2: Typical results for the BioTherm/RTI and WFBB processes (after Piskorz et al., 1998).

Yields, wt %, hardwoods sawdust feed Process Temperature,

°C (°F) Char (solids)

Bio-oil (liquid) Gas

WFBB 425 (797) 30.5 59.6 5.9 WFBB 500 (932) 7.7 78.0 10.8 ENSYN-RTP 425 (797) 30.5 59.6 5.9 BioTherm™ 430 (806) 12.5 74.3 10.1 BioTherm™ w/ char converter 430 (806) 4.5 73.0 19.0

Table 5.3: Pyrolysis results for polyvinyl chloride and polystyrene (after Scott et al., 1990).

Polystyrene PVC

Run 1 Run 2 Run 3 Temperature, °C (°F)

520 (970)

532 (990)

615 (1139)

708 (1306)

Amount fed, kg (lb) 0.020 (0.044)

0.013 (0.029)

0.018 (0.039)

0.020 (0.045)

Feed particle size, mm (in)

< 0.5 (0.02)

< 0.5 (0.02)

< 0.5 (0.02)

< 0.5 (0.02)

Yield, % feed char 9.1 ( n/a ) ( n/a ) ( n/a ) condensate 6.3 ( n/a ) ( n/a ) ( n/a ) HCl 56.0 ( n/a ) ( n/a ) ( n/a ) gases and losses 28.6 11.5 15.7 15.2 styrene ( n/a ) 76.2 72.3 75.6 other aromatics ( n/a ) 12.3 15.7 15.2

The basic process configuration consisted of the WFBB fluidized bed filled with Ottawa sand (250 – 590 µm in diameter) and fluidized with nitrogen gas. Gases and condensable vapours were collected upon exiting the reactor using a system of ice-water condensers, cotton-wool filters, an activated carbon adsorber, and a gas bag. Entrained solids were separated using a cyclone and collected in a char-pot. All components were weighed and analyzed using a gas chromatograph.

Typically, over 80% conversion into gaseous and liquid hydrocarbons was achieved when feeding polyethylene and over 70% conversion of polystyrene to styrene. The conversion of poly(vinyl chloride) PVC yields mostly HCl and a complex mixture of various other gases (Scott et al., 1990). Table 5.3 and Table 5.4 summarize some of the results found for the pyrolysis of poly(vinyl chloride), polystyrene, and polyethylene.

5.3 Ensyn – RTP

5.3.1 Introduction

The rapid thermal processing, or RTP, process developed by Ensyn originated from research at The University of Western Ontario (see Section 2.2). As stated, the original development was towards a pyrolysis system for the production of liquid and gaseous fuels from biomass (Graham et al., 1986). Later work investigated the production of light olefins including ethylene and propylene. The process was first commercialized in 1989 when a plant was sold to the Red Arrow Company in Wisconsin Illinois for use in the production of food flavorings and additives. Now, there are eight operational plants in total ranging in production capacity from 10 kg/hr (22 lb/hr) R&D facilities to the 3, 25, and 30 dry t/d (2700, 23000, 27000 kg/d) units at Red Arrow. As of 2000, plans were being undertaken for the design and construction of a 350 dry t/d (320000 kg/d) plant (Bridgwater and Peacocke, 2000). Many of these units have been constructed in Europe, like the VTT unit in Finland and the ENEL facility in Italy.

Other development work is concentrating on hot vapour and liquid filtration in order to reduce the amount of solid char found in the liquid bio-oil, and the supply of bio-oil for testing in diesel engines. Currently, Ensyn is the testing of the RTP process for the upgrading of heavy oil. Some preliminary results from such testing are presented in Section 5.3.4. As stated in Section 2.2 heavy oil has already been tested on the laboratory scale version

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Table 5.4: Pyrolysis results for polyethylene (after Scott et al., 1990). Linear Low Density Polyethylene

sand bed activated carbon bed Run 1 Run 2 Run 3 Run 1 Run 2 Run 3 Temperature, °C (°F)

654 (1209)

790 (1454)

725 (1337)

595 (1103)

730 (1346)

785 (1445)

Feed particle size, mm (in)

< 0.5 (0.02)

< 0.5 (0.02)

0.85 – 1.19 (0.033 – 0.047)

Yields, % feed char 5.9 0.2 0.0 7.1 11.3 13.8 condensate 51.1 10.3 11.3 17.4 25.8 25.6 gases H2 nd nd Nd 0.8 2.0 1.3 CH4 2.8 6.4 5.4 3.8 8.4 12.8 C2H4 10.4 31.1 19.3 4.8 15.0 19.2 C2H6 1.7 2.7 2.0 5.9 4.8 5.1 C3H6 2.5 12.8 8.1 8.1 11.9 10.6 C3H8 0.1 0.8 0.6 6.2 2.2 1.2 C4 0.0 7.7 7.5 12.8 7.8 0.8 C5 + 1.7 0.7 0.3 29.3 0.2 3.4 Total 92.1 94.5 87.9 96.2 89.4 93.8

of what eventually became the Ensyn RTP process. Ensyn has also recently entered into collaboration with Gulf Canada in regards to upgrading heavy hydrocarbons using the RTP process.

The heart of the RTP system is a transported bed reactor that contacts an inert particulate heat carrier, typically silica sand, with the incoming feed in an upflow design. The typical ratio of inert heat carrier to feedstock is between 20:1 and 30:1 but may range anywhere from 10:1 upwards of 200:1. The heating rate in the mixer is on the order of 1000°C/s (1832°F). The overall reactor temperature is generally less than 750°C (1382°F).

There currently exist four different configurations of the same basic system. In all configurations and for all types of feeds, the reactor residence time is generally on the order of a few hundreds of milliseconds, controllable upwards to a maximum of 5 s by relocating the feed injection point. The longer residence times are employed for the production of liquid fuels from biomass. For a wood biomass feed, the liquid yield is of the order of 83-wt %.

5.3.2 Process Description

Figure 5.7 shows a simplified process diagram for the RTP-III process of the ENEL Bastardo plant in Italy and the basic system components. Here biomass is fed into the reactor via a screw feeder – although different feed configurations are available for liquid or gaseous feeds. The product vapour is sent through two cyclones where any solids are separated before passing through two standard condensers. A portion of the non-condensable product is sent back to be used as fuel in the sand heater and fluidization gas in the reactor. The sand particles are continuously recirculated between the reactor and the sand heater. Typically, the first cyclone is a reverse flow cyclone (Freel and Graham, 1998) and the second (not shown) a standard high-efficiency cyclone although any suitable separation means could be substituted.

Figure 5.8 and Figure 5.9 show two different possible configurations of the mixing section located in the lower portion of the reactor. As shown, there may be one or several injection points. The average residence time of the various components in the mixing zone is between 0.005 and 0.030 s. The biomass throughput achievable is on the order of 2.3 – 3.4 kg/(m2.hr) 1700 – 2500 lb/(hr.ft2) of reactor cross-section. Within the mixer section a dense turbulent phase exists as compared to the upflow transported bed reactor where there is dilute phase. This maximizes the interaction between the heat carrier and feedstock thereby enhancing the heat transfer and mixing phenomena. Feeding may be accomplished using a mechanical screw (e.g. Figure 5.8) in the case of a solid feedstock or by injection in the gas of a liquid or gaseous feedstock (e.g. Figure 5.9). The entrance velocity, and hence the penetration, of the feedstock is increased with the use of an inert gas, which also prevents backflow. Typically, the gas is the recycled process gas that has already been cooled in order to prevent premature pyrolysis of

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Biomass feed

Nitrogen

Reactor

Propane Air

Sand

Sand heater

Product oil

Filter

Filter

Demister

Exhaust

Air Water Filter

Condensers

Figure 5.7: The RTP-III process flow plan (after Bridgwater and Peacocke, 2000).

Recirculation gas

Feedstock Flapper valve

Figure 5.8: RTP mixing section (after Freel and Graham, 1998).

Figure 5.9: Alternative RTP mixing section (after Freel and Graham, 1998).

the feedstock. In the embodiment of the mixing section shown in Figure 5.9 the solid heat carrier is introduced via the angled inlets that are focused on the central stream of feedstock at an angle of 60°.

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As mentioned above, there exist several reactor configurations, which are illustrated in Figure 5.10 (a) through (d) below. All four configurations show a mechanical screw solid feed mechanism. Figure 5.10 (a) and (c) illustrate the inclusion of constriction zones between the mixing section and the reaction portion, which is claimed to further enhance the mixing and heat transfer as well as reduce the residence time of the system. Figure 5.10 (b) illustrates a typical RTP configuration for when solids are being recirculated between the reactor and a combustion chamber (where the excess char is burned off in order to reheat the solid heat carrier particles). Here, the constriction is at the point of solids re-entry in order to accelerate the particles prior to mixing with the feedstock. Total residence time in the mixer and reactor combined is between 0.3 and 1.8 s. Finally, the exit from the reactor is placed according to the desired residence time as determined by the pressure balance and flooding conditions of the separation/recirculation system.

Recirculation gas

Grid plate

Screw feeder

Recirculation gas

Feedstock inlet

Recirculation gas

Screw feeder

Product stream

Grid plate

(a) (b)

Product stream

Recirculation gas

Grid plate

Recirculation gas

Feedstock inlet

Screw feeder

Reverse flow cyclone

Solids recirculation line with flapper valve

Product stream

Recirculation gas

Grid plate

Recirculation gas

Feedstock inlet

Screw feeder

(c) (d)

Figure 5.10: RTP reactor configurations (after Freel and Graham, 1999).

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5.3.3 Biomass Pyrolysis Results

Table 5.5 shows a set of results from three different series of tests conducted using an upflow transported bed recirculating reactor with an inner diameter of 0.15 m (6 in) and a 30 kg/hr (66 lb/hr) processing rate.

One potential criticism of circulating fluid beds is the complex hydrodynamics of such systems relative to other systems. Other criticisms include a higher solids attrition rate, incomplete solid separation by the cyclones, and that adequate heat transfer exists at large scale is yet to be proven (Bridgwater, 2001).

Table 5.5: Typical results from the RTP process for biomass feedstocks (after Freel and Graham, 1998).

Physical properties Typical values Feedstock hardwood hardwood lignin C, wt % 48.5 48.5 63.32 H, wt % 6.2 6.2 6.08 O, wt % 44.2 44.2 33.35 ash, wt % 0.6 0.6 ( – ) temperature, °C (°F) 520 (968) 500 (932) 550 (1022) residence time, s 0.69 1.4 0.8 Product Yields, wt % total liquid 72.5 70.0 55 gas 13.0 19.0 10 solid 14.0 11.0 35 C, wt % 56.4 56.4 67.0 H, wt % 5.6 5.6 6.9 N, wt % 0.17 0.17 0.1 S, wt % 0.0005 0.0005 ( – ) ash, wt % 0.15 0.15 ( – ) viscosity, cP at 40°C (104°F) 51 51 ( – )

5.3.4 Heavy Hydrocarbon Processing

As stated above, testing has begun using the Ensyn RTP reactor technology for rapid thermal processing of heavy hydrocarbon liquids. Suggested feeds include heavy oil, bitumen, atmospheric tar bottoms, vacuum tar bottoms, coal oils, residual oils, tar sands, shale oil, and asphaltic fractions (Freel and Graham, 2002). The reactor configurations used for this study correspond to those already described above along with the necessary modifications to the feed injection system. For this study, the feed was generally preheated to 80°C (176°F) in a surge tank using band heaters while all transfers lines were maintained at around 150°C (302°F).

For the study in question, several different reactor configurations were evaluated including single stage configurations with and without recycle as well as 2-stage and multistage designs. The recycle and multistage designs involve recovering the heavier portion of the liquid product and in the former case recycling it back to the original reactor while in the latter case it is sent through one or more separate downstream pyrolysis units. Generally, the reactor residence times are between 0.35 and 0.7s and the solids loading on the order of 30:1. For the multistage designs the initial pyrolysis unit was operated at a lower temperature (e.g. 510 – 530°C or 950 – 986°F) while subsequent units were maintained at higher temperatures (e.g. 590 - 750°C or 1094 – 1382°F) in order to avoid over-cracking of the initial light components formed.

Table 5.6 shows the properties for two example hydrocarbon feeds: heavy oil from Saskatchewan and bitumen from the Athabasca tar sands in Alberta. Table 5.7 through Table 5.10 summarizes a portion of the results obtained using the feeds given in Table 5.6 for the various reactor configurations discussed above. To begin, Table 5.7 includes a typical set of results and reactor conditions for single stage operation. Next, Table 5.8 summarizes a representative set of product distillation analyses for both single and multistage reactor configurations. Then, Table 5.9 compares the effect of recycle on the API gravity and liquid yield. Finally, Table 5.10 shows an example of the effect of using a 2-stage configuration on the product API gravity, density, viscosity, and liquid yield.

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Table 5.6: Feedstock characteristics (after Freel and Graham, 2002). Component Saskatchewan Heavy Oil

(SHO) Athabasca Bitumen

(AB) C, wt % 84.27 83.31 H, wt % 10.51 10.31 N, wt % < 0.5 < 0.5 S, wt % 3.6 4.8 ash, wt % 0.02 0.02 V, ppm 127 204 Ni, ppm Nd 82 H2O, wt % 0.8 0.19 Gravity API° 11.0 8.6 Viscosity, cSt at 40°C (104°F 6343 30380 at 60°C (140°F) 892.8 1268.0 at 80°C (176°F) 243.4 593.0 Aromatics, C13 NMR 0.31 0.35 Conradson carbon, wt % 12 ( - )

Table 5.7: Single stage reactor conditions and product characterization (after Freel and Graham, 2002).

SHO AB Temperature (°C) 530- 620 520 – 590 Heat carrier to feedstock ratio 20:1 to 30:1 20:1 to 30:1 Residence time (s) 0.35 – 0.7 0.35 – 1.2 Liquid: Yield (wt %) 71.5 - 84 74.6 – 82.7 API gravity 13.3 – 16.6 11.6 – 13.0 Density (15° g/ml) 0.962 – 0.977 0.979 – 0.988 Viscosity (cSt) at 40°C 15 – 300 25.6 – 205 V (ppm) 40 – 130 74 – 88 Ni (ppm) 10 – 50 24 – 27 Gas (wt %): total 7.2 – 11.8 N/A ethylene 16.6 – 27.0 N/A ethane 8.2 – 16.4 N/A propylene 15.4 – 30.0 N/A methane 21.0 – 24.0 N/A Conradson carbon (wt %) 6.6 N/A

Table 5.8: Product simulated distillation curves for Saskatchewan Heavy Oil (SHO) and Athabasca Bitumen (AB) (after Freel and Graham, 2002).

Single stage: SHO

TR = 538°C (1000°F) (wt %)

Single stage: (AB)

(TR = 538°C (1000°F) (wt %)

Multi-stage SHO: T = 543°C

(1009°F) AB: T = 538°C (1000°F) (wt %)

Fraction Temperature range, °C (°F)

Feedstock Product Feedstock Product SHO AB vacuum resid > 538 (1000) 51.5 37.9 51.5 32.0 12.7 19.5 pre vacuum resid < 538 (1000) 48.5 61.9 48.5 68.0 87.2 80.5

naphtha/ pre-kerosene < 193 (379) 0.1 3.2 0.1 2.7 0.7 6.2

kerosene 193 – 232 (450) 1.0 2.8 1.0 2.6 2.8 4.4 diesel 232 – 327 (621) 8.7 14.2 8.7 14.1 21.8 19.7 light VGO 327 – 360 (680) 5.2 6.5 5.2 7.3 10.8 9.1 heavy VGO 360 – 538 (1000) 33.5 35.2 33.5 41.3 51.1 41.1

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In general, as the reactor temperature is decreased there is an increase in the API, liquid yield, and viscosity, while the density decreases. There is claimed to be no significant product degradation observed, even after 30 days. The liquid yields obtained are typically on the order of 75 – 80%.

Table 5.9: Product characteristics from pyrolysis with recycle (after Freel and Graham, 2002).

Feedstock Temperature °C (°F)

Yield (vol. %)

API° Recycle Yield (vol. %)

Recycle API°

Heavy oil 590 (1094) 77.1 13.3 68.6 17.1 560 (1040) 86.3 16.2 78.1 21.1 550 (1040) 50.1 14.0 71.6 17.8 550 (1022) 65.1 18.3 56.4 22.9 530 (986) 87.1 16.6 78.9 21.0 Bitumen 590 (1094) 75.2 12.4 67.0 16.0

5.4 National Renewable Energy Laboratory

5.4.1 Introduction

The National Renewable Energy Laboratory (NREL) is a division of Midwest Research Institute (MRI) located in Kansas City, Mo. NREL has developed and patented an ablative vortex reactor system for the rapid pyrolysis of biomass for the production of bio-oils. Since the heat required for pyrolysis is transferred through the cylindrical vessel walls, no particulate heat carrier or hot gases are required.

Table 5.10: Two stage and multistage product properties (after Freel and Graham, 2002). Feedstock Temperature,

°C (°F) Yield (vol. %)

API°

Density (g/ml at 15°C)

Viscosity (cSt at 80°C)

515 (959) 31.4 18.6 0.943 5.3 590 (1094) 78.1 11.4 0.990 52.6

Heavy oil (2 stage)

Final 86.6 13.9 ( - ) ( - ) Multistage 510 – 750

(950 - 1382) ( - ) 15.9 0.9592 79.7

5.4.2 Process Description

The basic reactor configuration for the NREL design is given in Figure 5.11. As may be seen, the reactor consists of a vertically aligned cylinder. The feed enters tangentially through the inlet at the top and the product stream leaves through the axially aligned outlet at the bottom. In the configuration shown there is also included a recycle loop for returning unreacted feed particles back to the reactor. Tracer studies have indicated that the reactor operates essentially under plug flow (Diebold and Scahill, 1988).

Carrier gas, in the form of steam, is injected horizontally at the junction between the solids feeder and the recycle loop. The feed particles are introduced by means of a screw conveyor. An ejector acts to create the necessary pressure difference downstream to entrain the solid particles and conduct them to the reactor chamber. The gaseous and particle streams follow a helical path along the outside reactor wall. It is considered by the designers that a thin film of bio-oil on the wall surface facilitates the particle motion. An interesting modification was to include raised ‘ribs’ that force the two streams to follow an even tighter helical path than that which would normally occur. This increases the amount surface area used for heat transfer that the particulate stream comes into contact with and increases the residence time to the minimum required. Another relevant process modification was the design and installation of a special erosion plate, designed so as to be easily replaced, installed at the point where the solid particles first impact the reactor wall.

Feed

Figure 5.11: The NREL

vortex reactor system (after Diebold and Scahill, 1995).

Typically the carrier gas enters at a temperature between 400 and 750°C (752 and 1382°F). The vortex reactor wall is heated to approximately 625°C (1157°F). Diebold and Scahill (1995) state their belief that the rapid heat transfer achieved using this ablative technique is through conduction across the thin layer of bio-oil mentioned above. Solids enter the recycle loop at about 500°C (932°F)

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and cool to below 400°C (752°F) by the time they reach the screw feeder.

5.4.3 Results

It is reported by Diebold and Scahill that the vortex reactor system, although still in a relatively early stage of development,

has the scale-up potential to reach 230000 kg/d (250 ton/d). This scale would require a reactor approximately 0.5 m (1.6 ft) in diameter and 9 to 12 m (30 to 40 ft) in length. Most early experimentation was conducted with the intention of completely gasifying the solid biomass. However, more recent studies have begun concentrated on recovering the primary vapour and condensing them to produce bio-oil. A typical set of results is given in Table 5.11 using similar conditions to those mentioned above but with the reactor aligned horizontally as it was in the initial configuration.

Table 5.11: Elemental balance for NREL vortex reactor system (after Diebold and Scahill, 1988).

Char Liquid Water Gas 7.5 76.8 11.7 4.0

10.5 73.1 12.8 4.1 12.7 69.0 14.0 4.3

5.5 Cyclonic Systems for Biomass Fast Pyrolysis

5.5.1 Introduction

Several cyclone reactor systems exist for the fast pyrolysis of biomass for the production of bio-oils. The main body of work in this field appears to be from Lédé and colleagues from both CNRS-ENSIC-INPL (Lédé, 2000) and ELF Aquitaine (Soulignac et al., 1986 a, b). One further example is provided from Brem (2001). Tracer studies indicate a solid flow regime that is essentially plug flow and reactor residence times of the order of 1 s or less (Lédé, 2000; Soulignac et al., 1986).

The typical feed mentioned is that of wood sawdust, however it is stated that the feeds may include solid particles or liquid droplets (Brem, 2001) or heavy hydrocarbon feeds (Soulignac et al., 1986a). Soulignac et al.(1986a) also mention the importance of supplying uniform heat distribution over the entire surface of such cyclones so that the entire feed is subjected to the exact same conditions – a point that is applicable to all ablative reactors.

5.5.2 Lédé et al.

The studies conducted by Lédé and his colleagues have principally incorporated sawdust as the feedstock of choice in bench-scale cyclones. However, it would be possible to envision the direct utilization of tar sand as described in Section 2.5. The goal of their research was to confirm the behaviour of such systems under a wide variety of operating conditions in order to establish the necessary data required for scale-up. In their opinion there currently lacks the information needed for “reliable” scale-up and that most models are too empirical and apply only to the device with which they were derived.

Wood saw dust Ar

High temperature fan

Heating

Screw feeder Heating

Cyclone reactor

CondensersFilterFlowmeterGC analysis

Diaphragm

(recycling flow rate)

Figure 5.12: Experimental flow plan for the

studies by Lédé et al., (2000).

The cyclones used for the study were 0.028 and 0.040 m (1.1 and 1.6 in) in diameter (Lédé, 2000). Particles sizes ranged from 0.0002 to 0.001 m (0.008 to 0.04 in) in diameter with residence times of between 0.3 and 1.8 s. The gas residence time was typically between 0.08 and 0.15 s based on conditions at the inlet. The cyclone wall temperature was varied between 620 and 917°C (1148 and 1632°F) using either a 6-kJ/s (1.4kcal/s) solar furnace or an electrical heating element. A calibrated screw feeder was used to deliver the wood sawdust particles where it was mixed with a small amount of inert gas and steam. A flow plan for the experimental setup is given in Figure 5.12. Solid particles were recovered and removed via a small hopper located beneath the cyclone. The gaseous mixture leaves the

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(a) (b)

Figure 5.13: The cyclone pyrolysis reactor from Soulignac et al. (1986a) showing (a) side view and (b) top view.

cyclone at a temperature of between 550 and 700°C (1022 and 1292°F) and is sent through a condenser to separate the liquid products and steam while the aerosols then pass through a bed of cotton wool. A volumetric flow meter measures the total volume and instantaneous flow rate of gas while the composition is analyzed using gas chromatography. The entire process operates at atmospheric pressure. Both cyclones exhibited a maximum throughput depending upon the size of the cyclone and the operating temperature of 0.35 and 0.50 kg/h (0.77 and 1.1 lb/hr) at 927 and 1027°C (1700 and 1880°F) for cyclone A and B, respectively. A detailed view of the reactor configuration is given in Figure 5.13. It is interesting to note the solids inlet (indicated by the down arrow) that enters axially through the top of the reactor before entering an essentially horizontally aligned toroidal shaped section (refer to Figure 5.13 (b)). After a 180° bend the particulate feed particles are ejected tangentially into the upper portion of the cyclone. No mention is made as to the purpose of the horizontal conduits located on either side

of the feed inlet and gas outlet. However, Choi (1978) suggested for a similar process, the use of inert heat carrier particles or catalytic solids (see Section 6.1.3). Average results for a series of 22 experiments conducted by Lédé et al. are summarized in Table 5.12. The representative dry gas composition is given in Table 5.13.

Table 5.12: Typical mass balance results for the study by Lédé (2000).

Cyclone A Cyclone B Inlet temperature, °C (°F) 181 (358) 400 (752) Wall temperature, °C (°F) 1048 (1918) 893 (1639) Outlet temperature, °C (°F) 500 (932) 623 (1153) Duration of steam injection, s 5900 10800 Duration of pyrolysis, s 4304 5280 Reactants, kg x103 (lb x103) dry wood 344 (759) 290 (639) steam 1315 (2900) 4349 (9588) wood humidity 46 (101) 50 (110) total 1706 (3760) 4789

(10558) Products, kg x103 (lb x103) dry gases 279 (615) 309 (681) gas humidity 9 (19) 20 (44) solids 14 (31) 19 (41) condensed water and gases 1409 (3106) 4438 (9784) aerosols 4 (10) 10 (22) total 1715 (3781) 4796

(10573) Mass balance, % 100.54 100.14 Mass fractions, % dry gases 78.93 77.97 pyrolysis liquids 15.92 14.80 solids 3.93 4.71 aerosols 1.22 2.52

Table 5.13: Representative dry gas composition for the study by Lédé (2000).

Gas (vol. %) Cyclone A Cyclone B H2 27.9 21.7 CO 42.6 44.5 CO2 10.5 16.2 CH4 10.6 8.3 C2H2 1.1 5.3 C2H4 6.3 4.0 C2H6 0.8 4.0 C3H6 0.2 ( – ) Density, kg/m3 (lb/ft3) STP

0.946 (0.059)

1.100 (0.070)

Heating value, kJ/m3 (cal/ft3) x105

0.185 (1.25)

0.176 (1.19)

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5.5.3 Brem

Similar to the cyclone reactor systems for fast pyrolysis described above, the process described by Brem (2001) operates at close to atmospheric pressure with wall temperatures ranging between 500 and 800°C (932 and 1472°F) and a solids and gas residence time on the order of 500 ms. One major difference is that the cyclone reactor from Brem (2001) offers an enhanced filtering mechanism located directly in the cyclone exit. The filter mechanism consists of a housing containing several dozen chambers orientated parallel to the direction of flow of the gas exiting the reactor. These chambers may range in size anywhere from 1 to several millimeters (0.04 in) in diameter. Rotation of this filter housing along this same axis at speeds up to 3000 rpm creates a centrifugal force within each chamber forcing the entrained solids to be collected at the inner walls of each chamber. It is reported that such a system is capable of collecting particles anywhere between 0.1 and 50 µm. As the particles are collected they will tend to agglomerate. These particles are removed using a directed flow of inert gas and blown back into the cyclone, where they are more easily collected as a result of their increased size.

An illustration of the rotating filter is given in Figure 5.14. Figure 5.14 (a) shows a cross-section of several different possible arrangements of channeling within the filter, while Figure 5.14 (b) and (c) shows a side view of the filter mechanism and a close-up of a typical set of channels, respectively. An overview of the cyclone reactor configuration is given in Figure 5.15. Continuous operation has been achieved with this reactor system and filter

mechanisms upwards of 1 m (3.3 ft) in diameter are possible with equivalent gas flow rates of up to 10000 m3/hr (98 ft3/hr). The preferred embodiment of the reactor housing is in the form of a monolith made out of ceramic material

(b)

(a) (c) Figure 5.14: Rotating filter mechanism (after Brouwers, 1991).

The flow plan for the process described by Brem is shown in Figure 5.16. Allowance is made in the design of the cyclone for the inclusion of several inlets – i.e. one for the particulate feed and one for an inert heat carrier or catalyst. As may be seen from the process flow plan, the gaseous products exiting the cyclone are first quenched and then sent to a demister, which itself may include another rotating filter system. In fact, quenching may occur within this second cyclone. Finally, as with many other systems, the solid particles recovered are recycled back to the process via a fluidized bed heater where the excess char is burned in order to generate some of the heat required for the process. The remaining heat required would come from the heating of the vessel walls by any practical means.

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Two different reactor configurations are presented by Brem (2001) and are illustrated in Figure 5.17. For both designs the solids and carrier gas are introduced at the bottom of the cyclone with enough momentum so that the particle motion is upwards. The configuration shown in Figure 5.17 (a) includes an inner cyclone for collection and removal of solid particles while Figure 5.17 (b) shows a solids outlet located tangentially near the top of the reactor chamber. Finally, typical yields from this process are in line with those mentioned previously, namely: 80-wt % bio-oil, 20-wt % dry gases, and 10-wt % char. A set of typical results is presented in Table 5.14. It is not mentioned, but it is assumed that is for a feed of wood sawdust.

Carrier gas

Feed partic les

Reaction zone

Pyrolysis products

Rotating filter

Solid products

Flue gas

Ash

FB 850°C

Compressor

Pyrolysis reactor 550°C

Fluid jet

Quench

Rotating filter

Demister 50°C

Product gas

Bio oil

Cooling water

10°C

Recycle loop

Inert and char

Air

Dryer

Flue gas 100°CMill

Biomass

Figure 5.15: Cyclone reactor configuration

showing rotating filter housing (after Brem, 2001).

Figure 5.16: Flow plan for the pyrolysis process from Brem (2001).

Feed

Solids exit

Gas exit

Feed

Solids exit

Reactor

Rotating filter element

Gas outlet

(a) (b) Figure 5.17: Cyclone reactor configurations from Brem (2001).

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5.6 Shockwave Systems for Biomass Fast Pyrolysis

Shock wave pyrolysis systems have also been developed for the fast pyrolysis of biomass. One known example has been developed at Technische Universiteit Eindhoven (TUE). The pilot scale reactor system consisted of a reactor section 8 m (26 ft) in length and a shock wave generation system 4 m (13 ft) in length (Guo et al., 2002). The TUE system was tested using pulverized filtracel and lignin. Prior to injection of the biomass feedstock, the 4 m (13 ft) section is pressurized to 1.1 MPa (160 psia) using He gas.

When the diaphragm separating the two portions of the reactor is ruptured, the expanding gas creates a shock wave thereby raising the temperature above that required for pyrolysis.

Table 5.14: Representative results from Brem (2001). Bio-oil Moisture, wt % 15 – 30 Specific gravity, kg/m3 (lb/ft3) 1150 – 1250 (72 – 78) Heating value, MJ/kg (kcal/lb x102) 14 – 18 (1.52 – 1.95) Viscosity, cSt at 40°C (104°F) 30 – 60 Acidity, pH 5.5 – 4.0 C, wt % 50 – 60 H, wt % 5 – 7 N, wt % 0.1 – 0.4 S, wt % 0 – 0.1 O, wt % 34 – 43 ash, wt % 0.1 – 0.2

6 COAL PYROLYSIS PROCESSES

6.1 Occidental Petroleum

6.1.1 Introduction

The Occidental Research (Petroleum) Corporation (ORC) has been involved in the development of various processes for the pyrolysis of coal since 1969. Research work has involved both sub-bituminous coal and high volatile bituminous coal. Over the years several different design configurations have been patented, including a spouted bed reactor (Green, 1981), several types of coaxial jet reactors (Nuttal, 1978; Green et al., 1979), and a cyclonic reactor (Choi, 1978 a,b). The one commonality between all systems is the use of a particulate heat carrier, in this case solid char, in place of a gaseous heat carrier or direct heating.

6.1.2 Spouted bed design

Figure 6.1 (a) and (b) illustrate two basic configurations of the ORC spouted bed design for a pyrolysis reactor for agglomerative coals. As shown, the fluidized bed consists of hot char particles fluidized using any suitable non-oxidizing gas. Suitable gases listed include nitrogen, argon, methane, hydrogen, carbon monoxide, carbon dioxide, and steam (Green, 1981). Into this fluidized bed of hot char particles is introduced a central stream of finely pulverized coal particles at relatively high velocity (e.g. 24 m/s or 80 ft/s).

As the central jet expands it entrains the hot char particles and the pyrolysis process begins. Although of limited duration, it is at this stage that the risk of agglomeration is the greatest. However, the claim of Green (1981) is that the design of the reactor shown is such that by time the coal particles reach the vessel walls they are no longer adhesive and so do not agglomerate and plug the reactor. The inner chamber in Figure 6.1 (b) is designed to provide a certain protection against upsets in the char return line.

Referring to Figure 6.2, the gaseous products along with the entrained solids then exit the top of the reactor and are then sent to a series of cyclones for separation. The gaseous products are then rapidly quenched and the valuable liquid product sent for treatment. The remaining non-condensable gases are recycled for use in the char regenerator. The char particles recovered from the cyclones are split with a portion being returned to the pyrolysis chamber and the remaining being recovered as product. The coal particles used in this process are generally less than 32 mesh. The range of temperatures used for this process typically range from 590 to 870°C (1100 to 1600°F). Depending upon the heat balance considerations, the typical loading ratio of char to coal is 5:1 to 10:1. The solid char particles used in the process are no smaller than the coal particles themselves. In the injection stream, the solid coal represents no less than 10% and no more than 60% of the total jet volume. Factors that alleviate further any difficulties with agglomeration, include the use of a higher char to coal ratio, a higher reactor temperature, or simply a higher char temperature, and smaller coal particle size. In general, it is stated that the coal

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Hot fluidized char

Fluidizing gas

Vent

Cold coal feed

Hot fluidized char

Fluidizing gas

Vent

Cold coal feed

(a) (b)

Figure 6.1: Two alternative configurations of the ORC spouted bed design (after Green, 1981).

Air

Char Burner

Reactor

Cyclones

Cyclones

Combustion gas

Gas

Oil collection system

Liquid fuel

Liquid upgrading

Char product

Char desulfurization plant

Coal feed

Figure 6.2: ORC process flow plan (after Gavalas, 1982).

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particles are heated to the required pyrolysis temperature in less than 0.1 s (Green, 1981). No other reference is given to the residence time for the remaining system components, nor is any reference given to the relative product composition obtained.

6.1.3 Cyclonic Design

Figure 6.3 illustrates the basic configuration of the cyclonic concept developed at OCR by Choi (1978a). The potential feeds suggested include tar sands, oil shale, organic solid waste, coal, and mixtures thereof. Similar to a standard cyclone, the main components of the cyclone include the vertically aligned cylindrical portion, which merges downwards into a conical section to which is attached the coaxially aligned dipleg. Vapours exit from the upper portion of the cyclone. As seen in Figure 6.3 (b), the OCR design includes three separate entrances located side-by-side. The first (denoted by arrows), serves as the entrance for the heat carrier particles traveling at approximately 61 m/s (200 ft/s). The carbonaceous feedstock enters as the stream marked by “+” signs, while a third stream designated by “•” again contains heat carrier particles. The latter two streams enter at lower velocities relative to the first stream. In addition, it may be seen from Figure 6.3 (b) that the third stream enters at a marked angle relative to the first two. A non-oxidative carrier gas, if required, should be used to transport any of the given solid components. In some gases, reactive gases may be added in order to enhance the product obtained (e.g. hydrogen).

(b)

(a) (c)

Figure 6.3: Double helical cyclone pyrolysis system from OCR (after Choi, 1978a).

It is claimed that the inner feed of carbonaceous solids are kept separated from the wall of the cyclone by the faster moving solid heat carrier, which has a greater momentum (Choi, 1978). The pyrolysis time, counted from the moment of initial contact between the carbonaceous feed and heat carrier particles to the moment of separation, will be less than 3 s. Pyrolysis is carried out in this system at temperatures ranging from (480 to 760°C (900 to 1400°F), which maximizes the liquid yields. Similar to previous systems, the ratio of heat carrier to feed particle is from 2:1 to 20:1 (Choi, 1978b). Suitable heat carriers include sand or the solid char produced from the pyrolysis reaction. The feed velocity is generally between 31 and61 m/s (100 and 200 ft/s) – but no more than the “fast”

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stream of heat carrier. The distribution of particles between the two heat carrier streams is about 20-wt % for the “fast” stream and 80-wt % for the “slow” stream, respectively.

In most designs, the solid char formed on the heat carrier particles would be oxidized in a regenerator in order to produce the necessary heat for the endothermic reaction. The gaseous products and carrier gas are treated and separated by conventional means upon exiting the cyclone.

6.1.4 Coaxial Jet Design

Several variations of a coaxial jet design have been patented by ORC and are illustrated in Figure 6.4 through Figure 6.6. The primary differences between the three systems are the method and point of entry for the solid heat carrier

particles and the carbonaceous feedstock. In Figure 6.4 (a) the point of entry for the heat carrier particles is through an annular region axially aligned and surrounding the vertically aligned feed inlet. The solid heat carrier particles, in this case regenerated char particles derived from the pyrolysis process itself, and the heat carrier gas traverse a device shown in Figure 6.4 (b) consisting of several inclined vanes. These vanes act to induce a vortical motion to the heat carrier particles, which begin swirling around the outside of the cylindrical vessel. The stream of carbonaceous feedstock is injected into the centre of this swirling mass.

The main purpose of the swirling mass is to impede the agglomerative feedstock from reaching the vessel walls to quickly. This period allows the carbonaceous feedstock to be heated past the point where the surface is no longer adhesive. Baffles may be inserted at a strategic location in order to induce extremely turbulent mixing at the desired location and residence time

in the reactor. These baffles also act to stop the swirling motion of the outer stream prior to entry into the separation train.

Coal

Char

(a) (b) Figure 6.4: ORC coaxial pyrolysis system (after Nuttal, 1978).

Coal

Char

Char

Char

Char

Char

Char

Char

(a) (b) Figure 6.5: Coaxial jet designs from OCR (after Nuttal, 1978).

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Separation of the gaseous and solid components in this case is carried out using standard components. The condensable gaseous products are separated using condensers. The product char is combusted in an entrained bed heater before being recycled to the process. Variations on this method of introducing the heat carrier particles are shown in Figure 6.5 (a) and (b).

Fluidizing gas

Coal with carrier gas

Char

Weir

Figure 6.6: Coaxial pyrolysis system with weir

(after Green et al., 1979).

The final design presented is shown in Figure 6.6 below. In this design, the char particles are introduced in a fluidized annular region surrounding the feed inlet and separated by a weir. As the fluidized char particles overflow the weir they enter the mixing chamber and come into contact with the central turbulent jet of coal particles. The ratio of char to coal may range between 2:1 to 20:1 depending upon heat balance considerations, but is most usually about 10:1. Observed residence times range from 100 to 3000 ms (Berg, 1986).

Operational temperatures typically range from 500 to 800°C (932 to 1472°F), with a pressure of 136 kPa (5 psig).

Problems with this design though have been observed though and are reported in the literature (Gavalas, 1982). In particular, although the weir system was designed in order to act as a buffer to

Reactor temperature (deg.C.)

500 550 600 650 700 750 800

prod

uct y

ield

(wt %

DA

F)

0

20

40

60

80

100

gas 1.5 swater 1.5 swater 1.5 sgas 3 star 1.5 star 3 star 3 schar 3 s

Char

Gas

Water

Tar

Figure 6.7: Product yield versus temperature (after Gavalas, 1982).

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Residence time (s)

0 1 2 3 4 5

Tar Y

ield

(wt %

DA

F)

0

10

20

30

40

660 deg.C607 deg.C.

Figure 6.8: Tar yield versus residence time (after Gavalas, 1982).

upsets in the systems, fluctuations in the char flow rate were still observed to result in severe coking on the reactor wall and continuous was reduced to periods of less than 24 hours. Nevertheless, this system has been scaled up to a capacity of 38 kg/hr (84 lb/hr) in a process development unit.

Some typical results for this particular pyrolysis system are shown in Figure 6.7 and Figure 6.8. To begin, Figure 6.7 shows the product yields versus temperature for total residence times of 1.5 and 3.0 s (from the point of contact to quench) for the pyrolysis of subbituminous coal in a bench scale unit without recycle char. The maximum tar yield shown is about 20%. Similar efforts on the process development unit only achieved yields of 10% or less until it was discovered that replacing the N2 carrier gas with H2O or CO2 impeded the secondary cracking reactions on the char surface. Then the tar yields observed matched those of the bench scale unit.

Figure 6.8 then shows the evolution of the tar yield for increasing residence times for a bituminous coal. A maximum in the liquid tar yield of about 38% is observed at just under 1.0 s for an operating temperature of 660°C (1220°F). In this case, the type of carrier gas did not show any effect on the tar yield.

6.2 Rockwell International

During the 1980’s Rockwell International developed a flash hydropyrolysis process for converting coal into gaseous and liquid products that could be used as fuels (Ullman et al., 1986; Sinor, 1982). The reactor design is based upon the principle of a jet engine. Hydrogen gas is used both as a heat carrier and, when in excess, as a reactive gas. Figure 6.9 shows a general illustration of the horizontally aligned hydrogen heating coil and the vertically aligned mixing and reaction chamber of the Rockwell design. Located at the exit of the reaction chamber is the quench zone. Figure 6.10 shows a more detailed view of the injection and mixing zone.

The preheated hydrogen is injected through four circular ports at a relatively high flow rate of 0.0034 kg/s (0.0075 lb/s). These four jets are angled in such a manner as to impinge upon a central dense phase (i.e. close to its bulk density) stream of coal. The preheater must be specially designed for handling pressurized hydrogen. At the inlet, the heating coil has a relatively thin wall and small diameter but terminates with a relatively thick wall and

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large diameter. Seven motor generators are used to supply 600 – 800 amps that resistively heat the outer wall of the heating coil via two copper stud connects with a power input upwards of 150 kW (Sinor, 1982). The quench chamber consists of three rows with four full-cone spray nozzles that spray pressurized water.

(a)

Injection zone

(b) Figure 6.9: Rockwell design (after Sinor, 1982). Figure 6.10: (a) Top and (b) side views of

injection zone (after Sinor, 1982).

The reactor residence time is on the order of residence time on the order of 20 – 200 ms with a reaction temperature of 700 – 1000°C (1292 – 1832°F) (Bassi et al., 1993). The reactor pressure is maintained at 7 MPa (1015 psia). The mass ratio of hydrogen to coal is generally around 0.5. The overall coal throughput is on the order of 45 kg/(s.m2) or 33000 lb/(hr.ft2). The average particle size of coal used is between 40 and 50 µm. Generally the products consist of approximately 19-wt % synthetic crude with a boiling range of 200 – 350°C (392 – 662°F) and 9-wt. % dry gases including methane, ethane, and carbon oxides. Typically there is 3-wt. % of organic compounds recovered in the quench water. The current status of this technology, and the overall scale achieved, is not known.

7 CONCLUSIONS

As may be seen from this discussion, many different processes have been developed for fast pyrolysis. Obviously there is a strong economic incentive – to increase the yield of the more valuable liquid products – that is behind this effort. However, differences in design are also due in part to the various ideas on how to introduce the feedstock into the reactor. As well, differences have arisen concerning the mechanism of contacting the feedstock and the heat source as well as their subsequent separation. Still other small variations are evident in how the product vapours are quenched, and when, and what form the heat source takes, and finally, heat balance considerations.

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The preliminary results seen for the fast pyrolysis of heavy oils that were conducted at the laboratory scale are indeed economically attractive in that they do show a significant and positive increase in the liquid yields. Obviously, other important benefits would be derived from the use of such short contact time systems. Among other considerations, such systems would be more compact and hence their capital and possibly their operating costs would be lower. However, one study discussed here has shown that, at least as far as biomass pyrolysis is concerned, short residence times may not be the overriding principal concern. The good initial contacting that was present in all studies may play a crucial role: just as important as the overall residence time. In fact, it is argued by some that intense mixing of the two phases is more important than their uniform dispersion over the cross-sectional area of the reactor. Regardless, it is a necessity to achieve rapid and thorough initial mixing between the feedstock and the heat source.

Also, an important aspect related to the residence time, is the quick and efficient separation of the product vapour from the heat source. Not only is this essential in order to avoid degradation of the product due to secondary cracking but also it is crucial in providing the most efficient use of heat energy. The important aspect is to control the residence time in order to allow the reactions to proceed long enough to crack as much of the feedstock as possible and yet minimize the adverse secondary cracking reactions. This issue could be addressed by maintaining the residence time distributions as close to plug flow as possible, at least after the initial stage of mixing. Furthermore, it would appear from the literature that reactor systems operating in downflow regimes appear to be better to those that operate using an upflow regime. However, downflow systems may require more complicated flow systems to satisfy pressure balances and achieve a good seal between the reactor and the regenerator.

Finally, there is the fast separation of any particulate solids from the product vapours. Not only is this an important aspect for avoiding secondary cracking reactions as was already mentioned, but it plays a role in controlling the final product residence time as well. Besides, it is not practical to quench both the product vapours and the heat carrier solids (if any) since they would only need to be reheated anyways. One final point to consider would be to include a method for conducting relatively minor stripping of the solid material before it is sent for regeneration. Consider for example Figure 2.7 where the inclusion of several small stripping ports could potentially remove any product vapours trapped by the bed of particles that is formed. This would then increase the yield and minimize any secondary reactions occurring on the otherwise still hot particles, which would then minimize the coking and decrease the risk of fouling at this location.

Each process discussed has demonstrated some degree of merit in the form of an innovative technique for feedstock injection, mixing, separation, or quenching. However, each process inevitably has its critics and all are still under development. Perhaps then, the ideal process for fast pyrolysis would combine the best and proven features from each of these processes. Evidence supporting these arguments comes in the form that no one single process mentioned has been developed to the scale that would, for example, be required for Syncrude Canada Ltd., in order to process their current daily feed rate.

ACKNOWLEDGEMENTS

The authors gratefully acknowledge the financial support of the Natural Sciences and Engineering Research Council as well as Syncrude Canada Ltd.

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