December 2021

Process Optimization

Design and scale-up of gaseous Group A fluid bed systems for chemical synthesis

More than 100 fluid bed reactors for chemical synthesis and comparable processes using Group A powders have been installed and operated successfully since the late 1940s, with some reactors having an inside diameter at much larger than 7 m (23 ft).

More than 100 fluid bed reactors for chemical synthesis and comparable processes using Group A powders have been installed and operated successfully since the late 1940s, with some reactors having an inside diameter at much larger than 7 m (23 ft). This article discusses developmental milestones leading to their success, the key design differences between chemical synthesis circulating fluid beds and fluid catalytic cracking (FCC), and the author’s recipe for successful scale-up. Root causes of typical operational issues of this class of beds are also discussed.

Introduction

Types of gas fluid beds used typically include:

• Fixed fluid bed (FFB), also called bubbling or conventional fluid bed. Examples include phthalic anhydride, acrylonitrile, ethylene dichloride, maleic anhydride and Sasol gas-to-liquid (GTL).
• Dual FFB circulating bed. Examples includes Model IV FCC and others.
• Circulating fast fluid bed (CFFB), using a riser reactor. Examples include Sasol Synthol GTL, DuPont maleic anhydride, TRIG coal gasification, catalytic waste conversion, gas-to-olefin and chemical looping. They are characterized by a riser velocity of < 10 m/sec and very high riser solids holdup.
• Circulating fast bed (CFB), using a riser reactor. Examples include FCC, pyrolysis and coal combustion. They are characterized by a riser velocity of approximately 6 m/sec and low riser solids holdup.
• Moving beds and downers.

Fluid bed chemical synthesis reactions convert gas-phase hydrocarbons to high-value-added products. Reaction is typically limited to around 450°C (840°F). The more common mode of operation is with gas-phase oxygen supplied as air, enriched air or oxygen. When using gas-phase oxygen, operating pressure is usually limited to around 3 barg (45 psig) due to safety concerns and also because gas-phase oxygen reactions are favored by reduced pressure. Higher pressures may be used safely when oxygen is provided in solid form or when there is no oxygen, as occurs in some processes with comparable characteristics, such as the Sasol circulating or bubbling GTL process.

Since the mid-1940s, more than 100 FFB reactors using Group A powders have been installed and operated successfully for chemical synthesis and comparable processes around the world. Their success was based on the following developmental milestones:

• Correct overall range for catalyst size (wide size distribution, high levels of 0 micron–40 micron)
• Figuring out how to perform research at the lab to obtain data for scale-up correctly
• Correct fluidization regime(s) to use
• Proper gas distributor design (many holes distributed evenly with adequate pressure drop)
• Placement of coils inside the reactor and the process and mechanical design parameters of the coils.

Use of CFFB systems using Group A powders for chemical synthesis also dates back to the 1940s, but their use has been less successful than FFBs due to more complex design, higher risk when gaseous oxygen is present, and limited window of operation when internal coils are needed (metal erosion). CFFBs have been used or assessed for several types of large-scale processes:

• GTL by iron oxide reduction (solid-phase oxygen)—Sasol Synthol process
• Hydrocarbons to oxygenated chemicals by adsorbed oxygen (solid-phase oxygen)—DuPont butane to maleic anhydride
• Hydrocarbons to oxygenates with air, enriched air or oxygen (gas-phase oxygen)
• Natural gas to olefins by adsorbed oxygen (solid-phase oxygen)
• Natural gas to olefins with air (gas-phase oxygen)
• Non-slugging gasification of coal with oxygen (gas-phase oxygen)—TRIG process
• Conversion of solid waste to chemicals or syngas
• Chemical looping.

FIG. 1 depicts two CFFB systems presented in literature.¹ The sketch on the left represents the Sasol Synthol process for conversion of syngas to gasoline using iron oxide as catalyst. The iron oxide is reduced as conversion progresses. The sketch on the right is the concept piloted and commercialized by DuPont to convert butane to maleic anhydride using adsorbed oxygen. Both of these systems circulate a solid between two or three vessels, consisting of a riser reactor, a stripper/solid separator and, when required, a regenerator. The regenerator may operate as an FFB or a riser.

CFFBs have been assessed both in concept and through demo-scale production plants for the conversion of natural gas directly to olefins by most major gas field producers. So far, however, the economics have not justified further development—the main barrier being process yield.

CFFB vs FCC. A CFFB has many design aspects similar to that of an FCC, which is referred to here as a circulating fast bed (CFB). A Group A powder chemical synthesis CFFB will also contain a reactor that operates as a riser, but it should not be designed using FCC design guidelines for the following reasons:

• Use of oxygen gas in the reactor side of the CFFB poses safety considerations with respect to the design of the reactor in the event of loss of solids circulation. This concern is not present in FCC. Deflagration may occur in the CFFB reactor due to a drop in solids holdup caused by loss of solid circulation.
• Chemical synthesis catalyst may cost up to $110/kg ($50/lb) with a life of many years compared to less than $8/kg and a few months of life for FCC catalyst. In CFFB, the unit is typically designed to minimize attrition. In FCC, attrition of the catalyst is not a major design concern.
• The diameter ratio of regenerator to standpipe in modern FCCUs is very large. The ratio in a CFFB with the same size standpipe is significantly smaller as coke production is much less in CFFBs, or not present at all. Most regenerator standpipe entry designs used in FCC do not translate well to a CFFB.
• FCC uses atomized liquid feed in the reactor. A CFFB for chemical synthesis is likely to use gas feed, which creates challenges on how to mix the feed with circulated catalyst in the reactor bottom, especially if the reaction involves the use of oxygen (gas or solid phase).
• The FCC reactor is endothermic. The FCC regenerator is exothermic and may or may not require heat removal. The CFFB may require heat removal in both the reactor and the regenerator. The use of a catalyst cooler commonly used in modern, large-scale FCCUs is not practical in the CFFB reactor side, or often even for the regenerator.
• Fast separation of gas and solid at the exit of the riser in the FCCU is critical to minimizing gas yield. This separation is not as important in most, if not all, CFFBs.
• CFFBs require special considerations on how the reactor exit is connected to cyclones due to the high cost of the catalyst.
• Average density in the FCC riser is about 80 kg/m³ (5 lb/ft³). In the CFFB, riser density is typically more than 160 kg/m³ (10 lb/ft³). The CFFB reactor also operates at significantly lower velocity than the FCC riser. Moving such a dense mixture vertically up a tall riser at low velocity is not an easy task.
• FCC catalyst pores contain high-boiling-point components that are difficult to strip. CFFB catalysts typically does not have this issue, so stripping in the CFFB is generally easier.

Regenerator temperature control options

Approaches to controlling the temperature of the regenerator vary depending on the type of process being considered, ranging from external catalytic coolers for FCC to internal cooling coils or other approaches for chemical synthesis (TABLE 1).

Riser solid holdup

FIG. 2 may be used to estimate CFB riser solid holdup as a function of velocity and solid flux and applies to solid particle density of approximately 1,100 kg/m³–1,750 kg/m³. The vol% of solids can be obtained from the graph and multiplied by particle density. Commercial CFFBs typically operate with solids holdup above 10 vol% and fluxes over 300 kg/m²/sec.

Scale-up

The author’s experience suggests that there are no hydrodynamic reasons to limit the capacity of a fluid bed using Group A powder that has proper solid size distribution, operating velocity and design of internals. The factors that limit CFFB unit capacities (reactor size) are typically as follows:

• Plot space: The structure space (width, length and especially height) needed to connect three vessels (reactor, stripper, regenerator) together in a CFFB loop becomes excessively large as equipment size is increased.
• Cost factors: Shop fabrication is usually cheaper than field fabrication. Vessel diameters are, therefore, limited to about 4.5 m (15 ft) due to over-the-road accessibility with shop fabrication. Since FFBs have a single vessel, the structure height changes only a little with the size of the reactor. As a result, single-train field-fabricated FFB reactors of 7 m (23 ft) and larger are commonly used.

The pilot plant, demo plant or first commercial plant of a new process does not need to be optimal. The plant should work reliably and prove that the process concept is commercially viable. Optimization can be achieved by tweaking the second commercial plant. Admittedly, this approach may require some calculated risk. It is the faster approach to demonstrating a new Group A-based fluid bed process (CFFB, CFB or FFB). The alternative is to resolve all issues, thereby extending “time to market,” during which period engineering and research hours are spent increasing the “cost to market.”

The development of the Z-Sorb low-sulfur gasoline process is an excellent example of how collaboration between research and engineering—and taking some calculated risks—can expedite the time to market of a new process. The author was part of the team that designed, built and successfully demonstrated a 6,000-bpd unit within 16 mos of the project award, allowing ConocoPhillips to license the process before the short window of opportunity disappeared. Engineering started on Day 1, based on limited bench-scale data, and progressed while bench-scale research continued in parallel.

If the research at the bench scale is done correctly, then all the information needed to design a commercial, high-fines Group A reactor can be obtained. The only reasons for using a pilot plant is to build confidence (an admittedly valid reason) and to verify issues that cannot be addressed at the bench scale. Some of these issues may be addressed at much lower cost using cold mock-ups.

CFD capabilities in modeling gas/solids systems have increased tremendously and will continue to do so. CFD is a useful tool for supporting successful scale-up of new, large-scale fluid bed reactors using Group A powders. Unfortunately, some organizations rely on CFD results too heavily, without asking if the results make sense.

Many articles have been written on how solids behave inside a riser and how this may impact reactor yield and gas back-mixing. Many such articles were written years ago as bubbling beds were being introduced. Obviously, the design of the riser exit can cause recirculation of solids, and therefore gas. As for the rest of the riser, the higher the solids concentration of a riser, the more “well-behaved” it becomes. In addition, although chemical plants are designed to operate at full capacity, they sometimes must operate at reduced rates for extended periods of time. If a new process cannot handle turndown using riser technology, then perhaps the use of a riser is not the best choice?

The hydrocarbon processing industry witnesses an occasional paradigm shift that allows novel and untried approaches to become viable. However, it is also true that almost every process used to produce a chemical via the application of fluid beds has already been examined in the past, in some form or another.

Recipe for successful scale-up

A list of lessons learned for the successful commercialization of a new, large-scale FFB, CFB or CFFB process for chemical synthesis includes:

• Catalyst with acceptable characteristics is available (this is a given)
• Bench-scale research is carried out correctly; this is easier said than done. Many organizations do not obtain data correctly. The challenges we face are highlighted by the use of partial oxidation reactions represented by Eq. 1:

CxHy + O + (N, Cl, etc.) = Product + H₂O + COx                                                                     (1)

Partial oxidation reactions are favored by reduced pressure. The lower the operating pressure and the lower the concentration of hydrocarbon in the total feed, the higher the per-pass conversion and per-pass yield. Testing such a process using a low-pressure or an atmospheric pressure unit, and focusing on adjusting only the concentration of the hydrocarbon, masks this effect. This results in lower-than-expected yields when the unit is scaled up in the pressurized pilot/commercial plant.

Some partial oxidation reactions can also suffer from the formation of “color bodies.” These are trace byproduct(s) formed in ppm concentrations or less when the catalyst is not at its optimal condition. Color bodies must be removed, sometimes with great difficulties, to ensure that the end-user product meets the required color specifications (hence “color bodies”). The conditions under which color bodies are formed can be found only through testing at the lab or pilot plant.

FIG. 3 shows the impact of reactor size on yield for a high-fines Group A powder process. The data is masked, as it is proprietary, and is reported relative to the smallest unit tested. Constant pressure, temperature and gas contact time were maintained. FIG. 3 shows that yield remains essentially constant as size is increased, except at one scale, due to hydrodynamics caused by bubbles. Not all Group A powder processes show this behavior.

• Address issues immediately, rather than ignoring them or wishing them away.
• Take a total engineering approach to the design of the pilot plant and the first commercial plant. The new reactor always receives the most attention in the design phase, but it cannot operate if the front or back end of the process is designed or built poorly and is incapable of reliable operation.
• Engineering should collaborate with the research team as soon as the viability of the process is validated in the lab, and adequate and repeatable yield data is obtained to start the first-pass concept plant design. This exercise often leads to the discovery of gaps in research that can then be addressed before committing to the pilot plant. These gaps will eventually be discovered and will need to be addressed. Unfortunately, this discovery tends to happen after the engineering firm is hired and staffed up to design the pilot plant. Engineers charging hours will end up working at a less-than-optimal pace while waiting for the missing information, thereby increasing the project cost and schedule.
• Select the correct engineer for the pilot plant design. The local full-service or boutique firm may be a lower-cost option for the design and construction phase, but it may not be the correct choice. Remember, this is fluidization and solids processing. Despite the advancements in the field, it is still somewhat of a “black art.” If the technology developer is forced to spend money to fix problems before steady-state pilot testing can begin, then any perceived cost savings during the design and construction phase may be spent multiple times over. In addition, by this time, the process owner will have operating staff on the payroll just waiting around, which becomes expensive. Not every company has deep pockets to ride through such an unfortunate event or reassign idle operators to other activities.
• Correct process and mechanical engineering should be carried out for the pilot and the first commercial plant reactor, other fluidized vessels and their internals.
• Not spending the correct amount of money is another caution. There are reasons why the cost for 2-in. metal-seated full-port ball valves and an actuator range from $3,000–$25,000 (U.S. suppliers’ data only).

Common operational issues in chemical synthesis reactors

Common issues related to the operation in chemical synthesis reactors include:

• Poor yield/selectivity, caused by:
  • Reaction chemistry not studied correctly/insufficiently
  • Incorrect catalyst size distribution (too coarse/too fine)
  • Incorrect reactor velocity (too low)
  • Poor gas distributor design or damaged distributor
• Reactor instability, caused by:
  • Reactor velocity too high (high velocity can mean low bed density and runaway reactions)
  • Coil layout too restrictive (solid motion must be reasonably free to dissipate the heat generated)
  • Incorrect distribution of coils within reactor (coils must cover an adequate portion of the reactor)
• Catalyst attrition, caused by:
  • Catalyst is too soft
  • Incorrect gas distributor design (jet velocity too high)
  • Incorrect cyclone design (cyclone velocity too high)
  • Incorrect gas distributor fabrication/installation or insufficient quality check (either new or repaired)
• Catalyst deactivation, caused by incorrect grid plate design
• Erosion of internals, caused by incorrect placement of internals relative to gas jets
• Filter operation, caused by:
  • Design solid loading to filter too low
  • Design filter face velocity too high
  • Design filter cake density too high
  • Design gas inlet pipe layout incorrect
  • Blowback gas temperature too low, causing pore plugging.

Recommendations

Commercializing a new Group A fluid bed process requires extensive research to support catalyst development and prove reaction chemistry, but its commercial success ultimately relies on the capabilities of the engineers that scale up and design the first commercial unit for the process. As such, engineering provides a critical and complementary function to research in the development of any new multiphase process, of which fluidization is one of the more important types. HP

LITERATURE CITED

1. Jazayeri, B., Handbook of Fluidization and Fluid-Particle Systems, Chapter 16, Yang, W.-C., Ed., Marcel Dekker Inc., 2003.

The Author

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